Process Equipment and Plant Design: Principles and Practices [1 ed.] 9780128148853

Table of contents :
Process Equipment and Plant Design: Principles and Practices
Copyright
Dedication
About the Authors
Preface
Acknowledgement
Introduction
1 . General aspects of process design
1.1 Process
1.2 Design problem and its documentation
1.3 The design process
Qualitative considerations can be
Quantitative considerations
Optimum design
Design steps
1.3.1 Deliverables
1.4 Organisation of the Book
Further reading
Introduction
2 . Heat transfer processes in industrial scale
2.1 Introduction
2.2 Exchanger types
2.2.1 Recuperator
2.2.2 Regenerator
2.2.3 Fluidised bed exchanger
2.2.4 Direct contact heat exchanger
2.3 Flow arrangement
2.3.1 Countercurrent flow exchanger
2.3.2 Co-current flow/parallel flow exchanger
2.3.3 Cross-flow exchanger
2.3.4 Split flow exchanger
2.3.5 Divided flow exchanger
2.3.6 Multipass exchanger
2.4 Exchanger selection
2.5 Heat exchanger design methodology
Process and design specifications
2.6 Design overview for recuperators
2.6.1 Thermal design
The effectiveness-NTU method
2.7 Estimation of overall design heat transfer coefficient
Further reading
3 . Double pipe heat exchanger
3.1 Introduction
3.2 Design
3.2.1 Input data
3.2.2 Deliverables
3.2.3 Codes and standards
3.2.4 Guidelines to select inner and outer fluid
3.2.5 Design considerations
3.2.6 Thermal design
3.2.7 Hydraulic design
3.3 Series-parallel configuration of hairpins
3.4 Design illustration
3.4.1 Design steps
3.4.2 Design example
References
Further reading
4 . Shell and tube heat exchanger
4.1 Introduction
4.1.1 General description
Shell
Exchanger Head(s)
Tubes
Tube sheet
Baffle
Tie rods and spacers
Impingement baffle
Multipass exchanger
Shell passes
4.1.2 Heat exchanger installations and commissioning
4.2 Codes and standards
4.3 Design considerations
Process
Mechanical
4.3.1 Input data for design
4.3.2 Design output
Process design
Mechanical details
Fabrication details
4.4 Design – FT method
4.5 Pressure drop estimation
4.6 Mechanical detailing
4.6.1 Exchanger material
4.6.2 Tube length
4.6.3 Tube sheet details
4.6.4 Tube pass pattern
4.6.5 Finned tubes
4.6.6 Segmental baffles (transverse baffles in BIS code)
4.6.7 Tie rods
4.6.8 Impingement baffle
4.6.9 Shell dimensions
4.6.10 Channel and channel cover
4.6.11 Nozzles
4.6.12 Exchanger support
4.7 Design illustration
Further reading
5 . Heat exchanger network analysis
5.1 Introduction
5.2 Energy-capital trade-off – two-stream problem
5.3 Multi-stream problem
5.3.1 Optimal ΔTmin
5.3.2 Practical values of ΔTmin
5.4 Pinch design analysis
5.4.1 Locating the pinch using the problem table algorithm
5.4.2 The pinch principle
5.4.3 Design strategy
5.4.4 Grid diagram
Tick off heuristic
5.4.5 Stream splitting in network design
5.4.6 Network simplification: heat load loops and heat load paths
5.5 Targeting for multiple utilities
5.6 Design algorithm
5.7 Threshold problems
5.8 Data extraction
5.8.1 Composite curve for non-linear CP
5.8.2 Avoid mixing of streams at different temperatures
5.8.3 Use effective temperatures
5.8.4 True utility streams
5.9 Applications
5.10 Design illustration
Composite curves
Problem table algorithm
Further reading
6 . Evaporators
6.1 Introduction
6.2 Components of an evaporation system
6.3 Evaporator types
6.3.1 Types of continuous evaporators
Evaporators without heating surfaces
6.4 Evaporator performance
6.4.1 Multiple-effect evaporators
Feeding arrangements
Use of vapor as a “hot stream” in the plant
6.4.2 Vapor recompression
6.4.3 Heat recovery systems
6.4.4 Evaporator selection
6.5 Evaporator accessories
6.5.1 Condensers
6.5.2 Vent systems
Salt removal
6.6 Evaporator design
6.6.1 Single-effect evaporation
6.6.2 Multiple effect evaporation
Optimum number of effects in a multiple-effect system
6.6.3 Design data
Elevation of boiling point (BPE)
Boiling point elevation in multiple effect evaporators
Enthalpy plots
Tsteam & Tcon
Steam pressure
Pressure in the vapor space
Influence of feed, steam and condensate temperature
6.6.4 Design algorithm for multiple-effect evaporator
Design input
Design objective
Design deliverables
Design algorithm
6.7 Design illustration
Design example 1
Process design deliverables
Design example 2
Deliverables
Further reading
7 . Industrial cooling systems
7.1 Introduction
7.2 Cooling tower
7.2.1 Classification
Classification by build
Classification based on air draft
Classification based on airflow pattern
Classification based on the heat transfer method
7.2.2 Components of a typical cooling tower
7.2.3 Cooling tower parameters
7.2.4 Cooling water circuit in a process plant
7.2.5 Codes and standards
7.2.6 Thermal design
7.2.7 Notes on design and operation
7.3 Design illustration
Summary of available data
Tower selection
Fill details
Determination of operating L/G for the fill chosen
Steps of calculation
Fan power calculation
Estimating head loss in the fill and water distributor level
Estimating make up water (M) requirement
Evaporation loss (E)
Drift loss (D)
Pump calculations
Cooling tower sump
Further reading
Introduction
8 . Interphase mass transfer
8.1 Introduction
8.2 Processes and equipment
8.3 Process design and detailed design of the equipment
9 . Phase equilibria
9.1 Introduction
9.2 Representation of concentration
9.3 Representation of equilibrium
9.3.1 Graphical representation of equilibrium
9.3.2 Mathematical representation of equilibrium
VLE: Distillation
Solubility: absorption and stripping
GSE and LSE: adsorption
LLE: extraction
Further reading
10 . Absorption and stripping
10.1 Introduction
10.2 Tray column
10.2.1 Graphical determination of the number of contacting stages
Minimum required liquid flow rate (Lmin) in case of absorber for a given gas rate (G,G′)
Approximations for low concentration system
10.2.2 Absorption factor
10.3 Packed column
10.3.1 Packed column design based on mass transfer coefficient
10.3.2 Driving force line
10.3.3 Overall mass transfer coefficient
10.3.4 Estimation of active bed height
10.3.5 Design based on liquid-phase resistance
10.3.6 Absorption accompanied by chemical reaction
10.4 Design illustration
Driving force lines
Estimating mass transfer coefficients
Further reading
11 . Distillation
11.1 Introduction
11.2 Conceptual design
11.3 Detailed design
11.4 Fractionator
11.4.1 Process design of fractionating tower – equilibrium stage approach
11.4.2 Binary fractionation
11.4.3 Multicomponent distillation
11.5 Design illustration – fractionator
11.6 Flash distillation
11.6.1 Design equations
11.6.2 Design considerations
11.6.3 Design steps
11.7 Design illustration – flash distillation
11.8 Batch distillation
11.8.1 Design
11.8.2 Design deliverables
11.8.3 Design steps
11.9 Design illustration – batch distillation
Further reading
12 . Adsorption
12.1 Introduction
12.1.1 Modes of operation
Stagewise operation
Continuous contact operation
12.1.2 Adsorption mechanisms
12.1.3 Adsorption equilibrium
12.2 Packed bed adsorption
12.2.1 Breakthrough curve, breakthrough point, and bed exhaustion
12.2.2 Desorption/regeneration
Gas-phase adsorption
Liquid-phase adsorption
12.2.3 Adsorbent aging
12.2.4 Bed design
Rigorous methods
Empirical or short-cut methods
Pilot plant design
Data/information required for design
Operating parameters from pilot tests
(a) Loading rate/filtration rate (LR) for liquid-phase applications
(b) Superficial velocity (Us) for gas-phase applications
(c) Empty bed contact time
(d) Breakthrough time (tb)
(e) Fraction of bed utilised (f)
(f) Adsorbate loading (qs)
Bed design
Volume of fluid treated/change out period
Pressure drop
Bed configuration and mode of operation
12.3 Design illustration
Further reading
13 . Extraction
13.1 Introduction
13.2 Extractor types and selection
13.2.1 Extractor types
Stagewise contact
Continuous contact
13.2.2 Contactor selection
13.3 Choice of solvent
13.4 Design of continuous countercurrent contactors
Flooding
13.4.1 Calculation of the number of stages
13.4.2 Design parameters for extraction towers
13.5 Design of mixer-settler
13.5.1 Holding time
13.5.2 Power and mixing time
13.5.3 Scale-up
13.5.4 Flow mixers
13.6 Design illustrations
Further reading
14 . Column and column internals for gas–liquid and vapour–liquid contacting
14.1 Introduction
14.2 Tray towers
14.2.1 Contacting trays
Downcomer
Outlet weir
Liquid bypass baffles
Bottom tray seal pan
Weep holes
Vapour disperser elements
14.2.2 Choice of tray type
14.2.3 Tray construction
14.2.4 Efficient operation of contacting tray
14.3 Tray design
14.3.1 Bubble cap tray design
Tower diameter
Check for entrainment
Tray passes
Outlet weir
Height over weir
Downcomer area
Cap size
Number of caps
Area fractions over tray
Liquid gradient across tray
Tray pressure drop (htray, mm of liquid)
Check for vapour distribution
Vapour velocity and corrected ‘approach to flooding’
Downcomer pressure drop (hdc,prdrop, mm of liquid)
Downcomer backup (hL,dc, mm of liquid, for all cross-flow trays)
Velocity and residence time in downcomer
Downcomer throw over the weir
System (foaming) factors (applicable for all cross-flow trays)
Weep holes
14.3.2 Sieve tray design (cross-flow type – with downcomer)
Steps of design
14.3.3 Valve tray design
14.4 Packed tower
14.4.1 Choice of packing
Packing types and size
14.4.2 Liquid distribution
Liquid distributor
Redistributor and collector
14.4.3 Bed support
14.4.4 Flooding and pressure drop in randomly packed bed
Bed diameter estimation based on flooding and pressure drop
Pressure gradient
Minimum wetting rate
14.5 Packed tower design
14.6 Chimney tray, reflux entry, feed tray and tower bottom
14.6.1 Chimney tray
14.6.2 Reflux entry arrangement on top tray
14.6.3 Feed tray
14.6.4 Tower bottom arrangement
14.7 Design illustration
Further reading
Introduction
15 . Reactors and reactor design
15.1 Introduction
15.2 Design of reacting system
15.2.1 Reactor types
15.2.2 Rate and extent of reaction
Rate-limiting step
15.3 Reactor design
15.3.1 Reaction/process conditions
15.3.2 Design deliverables
Performance equation for idealized reactors
15.3.3 Scale-up
15.3.4 Bioreactors
Sterilization
15.4 Design illustration
Further reading
Introduction
16 . Plant hydraulics
16.1 Introduction
16.2 Pumps
16.2.1 Common pump types
Centrifugal Pump
Positive displacement pumps
Reciprocating pumps
Rotary pumps
Diaphragm pump
16.2.2 Pump performance and hydraulics
16.2.3 Cavitation
NPSH in centrifugal pump
Liquid vapour pressure
NPSH in reciprocating pumps
16.2.4 Characteristic curve for centrifugal pumps
Q-H curve
Pumps in series and parallel
Q-SHP (or BHP) Curve
Q-NPSHRCurve
16.2.5 System characteristic curve
16.2.6 Adjusting centrifugal pump performance
16.2.7 Characteristic curves for positive displacement pumps
16.2.8 Pump selection
16.2.9 Steps of design for a hydraulic circuit
16.3 Compressors
16.3.1 Compressor selection
16.3.2 Centrifugal compressor
Characteristic curve
16.3.3 Compressor hydraulics
Capacity and pressure ratio
Power
Head developed
16.3.4 Design/sizing
16.3.5 Capacity control
16.4 Piping
16.4.1 Piping codes
16.4.2 Pipe size
16.4.3 Piping services
16.4.4 Pipe rack
16.4.5 Pipe joints
16.4.6 Pipe fittings
Pressure relief–safety devices
Other fittings
16.4.7 Pressure drop in pipeline
16.4.8 Few typical process piping systems
Purge out operation
Vent and drain system
Flushing connections
Control valve installation
Steam trap
Good practices for piping layout
16.5 Hydraulic calculations
Further reading
17 . Process vessels
17.1 Unfired pressure vessels
17.2 Vessel components and fixtures
17.3 Mechanical design
17.3.1 Design Parameters
17.3.2 Vessel sizing
Vapour-liquid separator
Separator with wire mesh mist eliminator (demister pad)
Reflux drum
Liquid-liquid separator
17.3.3 Nozzle dimensions and location
17.3.4 Manhole specifications
17.3.5 Wall thickness
17.4 Design illustrations
Further reading
Introduction
18 . Utility services in process plants
18.1 Introduction
18.2 Fuel systems
18.2.1 Fuel gas
18.2.2 Fuel oil
18.2.3 Design of fuel system
18.3 Electrical power
18.4 Steam
18.5 Compressed air
18.5.1 Air supply scheme
18.5.2 Design illustration – compressed air system
18.6 Inert gases
18.7 Water
18.8 Efficient use of utilities
Further reading
19 . Plant instrumentation and control
19.1 Introduction
19.2 Control loop
19.2.1 Feeback and feedforward
Selection–feedback versus feedforward
19.2.2 Characteristic features of a process being controlled
19.3 Analog signals–pneumatic and electronic
19.4 Control algorithms
19.4.1 P, PI and PID controllers
Choice of P, PI, or PID controller
19.4.2 Few advanced configurations of controllers
Cascade control
Split range control
19.5 Measurement of process parameters
19.5.1 Temperature measurement
Thermocouple versus RTD
19.5.2 Pressure measurement
Measurement of differential pressure
19.5.3 Flow measurement
19.5.4 Level measurement
19.6 Control valves
19.6.1 Fail-open and fail-close valves
19.6.2 Valve size
19.7 Instrumentation for safety
19.8 Distributed control system (DCS)
19.9 Control schemes for common processes
19.9.1 Distillation control and instrumentation
19.9.2 CSTR instrumentation and control
Further reading
20 . Engineered safety
20.1 Introduction
20.2 Hazardous area classification
20.3 Trips and alarms
20.4 Blowdown and flare
20.4.1 Blowdown
20.4.2 Safety and pressure relief valves
20.4.3 Flare system
20.5 HAZOP
Problem statement
Report
Major recommendations
Worksheets
Worksheet WS–1
Worksheet WS–2
Further reading
Introduction
21 . Process packages
21.1 Process package deliverables
21.2 Examples
21.2.1 Design illustration 1
Design of 10,000 MT/Annum plant to manufacture Ethyl acetate from Ethanol
21.2.2 Design illustration 2
Design of a facility for a refinery to treat 8000m3/d of wastewater
Further reading
Graphical symbols for piping systems and plant
Based on BS 1553: PART 1: 1977
Scope
Appendix B: Corrosion chart
Physical property data bank
Conversion factors
Typical fouling factors in m2K/W compiled from various sources
Heat exchanger tube sizes and other details
List of different standards commonly used
Index
A
B
C
D
E
F
G
H
I
J
K
L
M
N
O
P
Q
R
S
T
U
V
W

Citation preview

Process Equipment and Plant Design Principles and Practices

Subhabrata Ray Department of Chemical Engineering, Indian Institute of Technology Kharagpur, Kharagpur, West Bengal, India

Gargi Das Department of Chemical Engineering, Indian Institute of Technology Kharagpur, Kharagpur, West Bengal, India

Elsevier Radarweg 29, PO Box 211, 1000 AE Amsterdam, Netherlands The Boulevard, Langford Lane, Kidlington, Oxford OX5 1GB, United Kingdom 50 Hampshire Street, 5th Floor, Cambridge, MA 02139, United States Copyright © 2020 Elsevier Inc. All rights reserved. No part of this publication may be reproduced or transmitted in any form or by any means, electronic or mechanical, including photocopying, recording, or any information storage and retrieval system, without permission in writing from the publisher. Details on how to seek permission, further information about the Publisher’s permissions policies and our arrangements with organizations such as the Copyright Clearance Center and the Copyright Licensing Agency, can be found at our website: www.elsevier.com/permissions. This book and the individual contributions contained in it are protected under copyright by the Publisher (other than as may be noted herein). Notices Knowledge and best practice in this field are constantly changing. As new research and experience broaden our understanding, changes in research methods, professional practices, or medical treatment may become necessary. Practitioners and researchers must always rely on their own experience and knowledge in evaluating and using any information, methods, compounds, or experiments described herein. In using such information or methods they should be mindful of their own safety and the safety of others, including parties for whom they have a professional responsibility. To the fullest extent of the law, neither the Publisher nor the authors, contributors, or editors, assume any liability for any injury and/or damage to persons or property as a matter of products liability, negligence or otherwise, or from any use or operation of any methods, products, instructions, or ideas contained in the material herein. Library of Congress Cataloging-in-Publication Data A catalog record for this book is available from the Library of Congress British Library Cataloguing-in-Publication Data A catalogue record for this book is available from the British Library ISBN: 978-0-12-814885-3 For information on all Elsevier publications visit our website at https://www.elsevier.com/books-and-journals

Publisher: Joe Hayton Acquisitions Editor: Anita Koch Editorial Project Manager: Lindsay Lawrence Production Project Manager: Paul Prasad Chandramohan Cover Designer: Matthew Limbert Typeset by TNQ Technologies

Dedicated to Our Students with Whom we have been enjoying a Rich Learning Experience over the years

About the Authors Subhabrata Ray Subhabrata did his graduation and post graduation from the Department of Chemical Engineering, Indian Institute of Technology Kharagpur, and has teaching in his Alma Mater since 1988. Prior to that, he worked with M/s Indian Oil Corporation Ltd. and M/s TECHNIP-ESIA in various capacities. Dr. Gargi Das Gargi had topped the undergraduate and postgraduate courses in chemical engineering from Jadavpur University and the Indian Institute of Technology Kharagpur, respectively. She is also the recipient of various awards for her research activities. She has been in teaching in the Department of Chemical Engineering, Indian Institute of Technology Kharagpur, since 2001.

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Preface A course on plant and equipment design would be incomplete without a perspective of the relevant process system. Thus, the design of a distillation column must include at least the broad details of the associated heat exchangers, piping, safety arrangements and instrumentation. The book is based on this approach e to include the auxiliaries of a system, be it any specific equipment or a process. Another specific feature of the book is a separate section on process utilities, instrumentation and controls and engineered safety of plants. Ideas about these systems, although essential, have traditionally received less prominence. The last section of the book discusses design of the complete plant e the process packages with design examples. It summarises the concepts and practices discussed in previous sections and demonstrates their applications in practice. This section is meant specifically to provide an idea of the capstone projects that the undergraduates may take up as part of their course curriculum. To expose the learner to the industrial standards and practices, the importance of relevant industrial codes and practices has been emphasised and referred to wherever necessary, albeit in a limited way. It is felt that the rookie engineers in industry may find this useful as a resource that covers the basic concepts along with industrial practices in processes and equipment designs. The way the book is structured, the user may start with Section I and cover Sections II, III and IV in any order. One may go through Section V as per one’s interest or as a part of a formal design course. Subhabrata Ray and Gargi Das Department of Chemical Engineering Indian Institute of Technology Kharagpur

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Acknowledgement This book bears the contributions from a number of individuals and institutions. These are everywhere even though only the names of the authors appear on the cover. Writing this volume was a pleasure but there were moments of uncertainties too that could be overcome only by the constant support and encouragement of our friends and colleagues. Though the names are numerous, we must mention Mr. Ashis Nag, Mr. Matilal Bhattacharya and Dr. Susmi Banerjea. They not only supported us but also contributed actively with technical information on current industrial practices. Their constructive criticisms time and again were invaluable. The constant encouragement received from our colleagues Prof. Sunando Dasgupta, Prof. Somenath Ganguly and several others kept us going. Throughout the period, Prof. B. C. Meikap and Prof. Sudipto Chakraborty supported us with their constant cooperation and pertinent suggestions. Many of our personal friends including Prof. Anand Patwardhan, Dr. Mohit Ray, Prof. Bitasta Chanda and Prof. G. Harikrishnan did not let us lose our focus in difficult times. Over the years of teaching, our students have expressed the need for a comprehensive reference text addressing different aspects of a process being designed. Their numerous feedback has resulted in this book. We proudly acknowledge their inquisitive expressions and are grateful for the same. Several of our students, now working in industry, guided us to shape the content with their ideas on the requirement from industry perspective. The names of Dr. Balaram Suman, Dr. Sahil Malhotra, Mr. Harshit Madan and Mr Abhishek P. are only a few in the endless list. The authors acknowledge Mr. Lawrence Lindsay, Ms. Anita Koch, Ms. Sheela Bernardine Josy, Mr. Chandramohan Paul Prasad and the team from M/s Elsevier who bore with us over the long years and acceded to our demanding requests. Without their support, publication of this book would never have been possible. The authors appreciate the support of Indian Institute of Technology for this wonderful experience. Subhabrata acknowledges his earlier employers M/s Indian Oil Corporation and M/s TECHNIP-ESIA for allowing him the opportunity of learning in industrial environment. Long hours of occupation by the authors working on the book have been borne patiently by our family members. Their patience and unstinted support is gratefully acknowledged. Subhabrata Ray and Gargi Das Department of Chemical Engineering Indian Institute of Technology Kharagpur

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SECTION

Introduction to process design

I

Successful design is not the achievement of perfection but the minimisation and accommodation of imperfection eHenry Petroski

CHAPTER

General aspects of process design

1

1.1 Process An industrial process converts feed material to useful product(s) of desired quality in commercial scale. The steps involved in such a process need to be technically viable, safe, and economical. Such steps involve heat, mass, momentum transfer, and chemical reaction, independently or in combination. A process is always designed to perform a specific function. Such a task can be heating of a material from an initial to a final temperature, mixing of several streams to achieve homogeneity, separation of a multicomponent material stream or chemical conversion of a reactor feed to products and their subsequent separation.

Simple and complex processes

Simple processes are usually centered around one single or major equipment. Auxiliary equipment may be required to complete the functionality. Examples of simple processes can be:

Process

Equipment

Auxiliary equipment

Heating of a stream

Heat exchanger

Pump, pipe fittings

Sieving for separation of solids

Sieve

Receiving of naphtha from berthed tankers to shore tanks and its transportation to a petrochemical plant a few kilometers away

Pump and pipeline

Storage tanks at berth, if required

Capture of hydrocarbon solvent vapor from a process by adsorption in an activated carbon bed

Packed bed

Blowers, piping, and pipe fittings

Separation of a mixture of benzene and toluene by distillation

Distillation column

Heat exchangers (reboiler, condenser), pump, and pipe fittings

Drying of compressed air for supply to instruments

Drier

Compressor, filter, pump, and pipe fittings, compressed air storage

Complex processes constitute of several simpler processes. The process plant converting raw material to product streams may involve reaction, separation by distillation, extraction, absorption, Process Equipment and Plant Design. https://doi.org/10.1016/B978-0-12-814885-3.00001-4 Copyright © 2020 Elsevier Inc. All rights reserved.

3

4

Chapter 1 General aspects of process design

drying, filtration, or for that matter any number of “unit operations” and/or “unit processes.” Such examples can be: • • • • • •

Production of common salt involving crystallization, evaporation, filtration, etc. Effluent treatment in plants employing several processes required for removal of specific pollutants present Coke oven plant that produces coke of different grades and by-products, starting from coal Catalytic hydro-desulfurization of naphtha that involves reactors and separation systems Air separation plant to produce oxygen and nitrogen involving compression, Joule-Thomson cooling, liquefaction and distillation Distillery consisting of several processes like fermentation, distillation, etc. to produce the saleable product “spirit.”

Large chemical complexes contain several interlinked plants/processes. Each plant produces stream(s) that is either a marketable product or is feed stream to another plant within the complex. There are multiple streams of raw materials, products and by-products within the complex. The “chemical complex” as a whole produces “products” that are marketable. As examples of large process complexes, one often talks about the Steel Plant, Fertilizer Plant, Petroleum Refinery, etc. For example, in a petroleum refinery, crude oil is fractionated to several component streams in its Crude Distillation Unit (CDU). Although the main operation, in this case, is “distillation,” the facility contains process steps, such as preheating of crude oil in a heat exchanger network with various hot streams, heating of crude oil in a furnace before it enters the distillation column, washing of some of the streams with caustic solution and subsequently with water to remove sulfurous impurities and so on. Several process plants like cracking units, hydrotreating unit, etc., in a refinery complex, further process the streams from the CDU. A process can be designed to operate in “batch mode” or in “continuous mode.” Batch processes for the same annual capacity will require bigger equipment but offer greater control on the process. Batch operations are preferred when the processing capacity is low, but the required control over process parameters is stricter. Also, when the product value is high compared to the investment in the plant Batch vs Continuous equipment, the economical option is often a batch process. These processes also offer greater variation and flexibility in the use of the equipment and its operation. As an example, a batch dryer for drying a specific product can be utilized for drying other products whenever it is free. Widespread applications of batch processes in the pharmaceutical industry dealing with relatively lower processing capacity and requiring stringent quality control, clearly illustrate these facts. In contrast, a continuous plant will require smaller equipment for the same annual capacity. This allows attainment of higher processing capacity for the same investment. Most large scale/high capacity industries like fertilizer, steel, petroleum refining, petrochemicals are continuous process plants as it turns out to be the more economic option. Some plants with low to moderate capacity use combinations of the batch, as well as continuous processes. In these plants, the output streams from the batch process are stored and processed later in the continuous processes. In some plants, the order of processing may be reversed. Use of continuous process avoids interruptions in processing that entail spending energy to start and shut down the process. Processing steps that take longer time or are difficult to control or require closer checks and controls operate in “batch mode.”

1.1 Process

5

Engineering design is not just a scientific solution. Merely achieving the functional goal is not sufficient. The solution must achieve its functional goal economically. Improved technical performance with the implementation of a “better” or more efficient design is always associated with the cost to be incurred. Design Therefore, the industrial process design has to be an optimum design, balancing the improved performance and increased cost. An example can be designing a heat exchanger to cool a stream of 70 m3/hr of kerosene in a refinery from 90 to 40 C. This is the design objective for the process and may be achieved by several options: • • • • •

use of a water-cooled shell and tube heat exchanger use of a water-cooled double pipe heat exchanger use of an air-cooled heat exchanger (fin fan cooler) direct contact with an immiscible cold fluid (water) followed by gravity separation in a settler. utilizing some specific cold process stream that needs to be heated; this can be an attractive design option as it would save energy

There may be many more exotic options for achieving the same design objective. The acceptable design alternative has to meet the design objective and strive to be economically optimum. In addition, the following are also considered: •

Practical deviations in design input parameters

It needs to be appreciated that the design inputs are associated with a certain level of uncertainty and the design solutions are worked out based on the best estimates of these inputs. Inputs may Design Considerations also change over time due to change in operating conditions of other sections in the plant, variation in market demand or product specifications and raw material quality. There can be other reasons too. For example, the kerosene stream known to be available at 90 C in the design stage of the heat exchanger may be available at a slightly higher temperature (say 94 C) after the equipment has been installed. A “robust design” needs to accommodate, up to a limit, such uncertainties particularly in quality (temperature, pressure, composition, etc.) and quantity (flow rates) of the input streams for the process. Safety factors or design margins take care of this during design calculations but leads to “overdesign.” This is illustrated in the design problems discussed later in the book. Nevertheless, the extent of overdesign needs to be optimal as costs increase with a margin of overdesign. Such limits in specific cases are decided based on industry practice, designers’ experience, and the criticality of the function of the designed process/ equipment. Typically heat exchangers are overdesigned to the extent of 10%e20% of their rated heat load. •

Compatibility with rest of the plant or the process complex

Compatibility is enforced by clearly defining the specifications, specifying the design codes and material standards. Such compatibility is warranted in terms of mechanical features, reliability and process considerations. Mechanical compatibility is essential for connectivity of different equipment. This is enforced by using the same standards for the flanges and fixtures for “hooking up” the designed equipment to existing or other equipment that may have been delivered from a different design group.

6

Chapter 1 General aspects of process design

•

Reliability of the process and mechanical design of equipment This is ensured by following the relevant standards and codes

• • • •

Available space for the equipment or process being designed Scope of interchangeability with other equipment or common spare Ease of operation and maintenance Safety during erection, commissioning, operation and maintenance.

The process designer needs to be absolutely clear about the functional requirements to be met by his design output and looks for the following minimal information at the outset: • •

•

•

Input and output stream detailsdrange of flow rate, temperature, pressure, composition or any other available information Required capacity of the processddefined in terms of maximum, normal and minimum flow rate (input and/or output) of streams. In some cases, this may even be stated in terms of energy to be supplied or removed from the system Spatial location and its limitations, the plot plan, and the preferred orientation of equipment. The information of the surrounding area, its geographical details, presence of any existing habitation around the proposed process plant, topology of the plot, details on wind direction (Wind rose), location of effluent discharge points, preferred dimension of standard equipment or its components like heat exchanger tube length, etc., are all to be considered in the design phase to deliver an optimum option. Specific standards to be followed for designing, e.g., IS 2825/ASME Sec VIII Div.1 for unfired pressure vessels or TEMAeR for heat exchangers. This is essential for generating a design that is compatible with the rest of the facility, use of standard material of construction and ensuring the reliability of the equipment. Standards greatly reduce the time to deliver a reliable design and get it fabricated. The role of design codes and standards during the design activity, therefore, need no further emphasis.

Based on the available information, the designer fixes the conditions at the boundary of the process. In industry parlance, this boundary is the “battery limit.” Battery area refers to the physical area within which the process resides. Input and output stream details are to be known at the “battery limit” before proceeding further.

1.2 Design problem and its documentation Design requirements emerge out of an expression of a functional need. Such a need can be expressed for a new plant or processing facility. The need for a plant to process naphtha available from a local refinery and convert it to some useful petrochemical product can be the example of such a functional need for the design of a petrochemical complex. It is also possible to have the “need” for some expansion/modification of existing processes/plants. The plant design gets completed by designing the various equipment of the plant and its piping and utilities in an integrated fashion. The earlier mentioned example of an arrangement to cool a continuous (70 m3/hr) stream of hot kerosene oil (from 90 to 40 C) before the same is sent to storage is a simple statement expressing the need for an additional process facility to offer the required functionality, i.e., cooling. It is needless to say that even

1.3 The design process

7

such a process system would include one or more equipment; possibly, a heat exchanger with allied auxiliary facilities like pumps, piping and pipe fittings that would serve the purpose. The “expression of functional interest” is followed by documenting the “preliminary design specifications” (PDS document). This goes beyond just the functional aspects of the proposed design and attempts to identify and quantify the design deliverables like capacity, material and process fluid properties, process stream conditions at inlet and outlet of the process, etc. In spite of containing additional information, the formulation of the design problem at this stage will hardly be complete in all aspects to embark on the design process. Generating the complete definition of the design problem is the next step and the outcome of this step is documented as “detailed design requirements” (DDR document). This document is prepared at the designer’s initiative with the aim of defining all aspects of the design problem before embarking on the design procedure. This is based on the response to a questionnaire set framed by the designer specific to his design assignment. This questionnaire set is usually called “Basic Engineering Design Questionnaire (BEDQ)” and requests information under various subheadings. Preparation of the BEDQ document is the joint responsibility of the client and the designer. This serves as the designers’ starting basis and is consulted by the designer as a reference in almost every step of his activity. BEDQ documents important issues, such as the brief background of the design case, design capacity, size/ dimensional limitations and applicable standards/codes for designs, drawings and materials. This document even specifies the format of the design deliverables and their documentation standards. Freezing the mutually agreed version of the “detailed requirements design document” (DRD document) marks the start of the design activity. Complete design of a process plant consists of the Basic Engineering, followed by Detail Engineering. Basic Engineering includes process selection, equipment selection, PFD/P&ID, functional description of the system and overall Plot Plan. The process design falls under Basic Engineering. Equipment mechanical design; piping drawing with support structures, etc., are covered under detail engineering. Basic Engineering for the process is usually supplied by the technology vendor in the form of a “Basic engineering package,” and the detailed engineering is carried out by a Detailed Engineering contractor.

1.3 The design process Designing always starts in a concept design phase, where the designer considers potential alternatives and compares those heuristically. Based on heuristics and experience, a finite number of alternatives are selected. For example, in the kerosene cooling problem, the evaluation of the options considered by heuristics will probably lead to employing a shell and tube exchanger over the other alternatives. Selection of design solution is usually based on qualitative considerations followed by quantitative considerations.

Qualitative considerations can be • • • •

Soundness of the scientific concept Fire and Safety risk of plant equipment, manpower involved and surrounding area Feasibility of practical implementationdthis includes compatibility with existing components of a plant, in case of retrofit or plant expansion, related designs as well as spatial limitations Economic attractiveness or advantage of one option over other(s)

8

• • • • • • •

Chapter 1 General aspects of process design

Requirement of auxiliary facilities like utilities and their levels Waste disposal/environmental considerations Ease of operation, maintenance, erection, commissioning Availability of resources for technology, materials, manpower and skill for erection, commissioning, operation, and maintenance Project completion time Designs with proven implementation and performance are preferred Financing of the project: Capital availabilitydamount and its layout in a time scale

In practice, the first step of the designer is the collection of information and previous documents on similar design projects. Implementation success, performance and limitations of these are scrutinized. The design alternatives emerging from this step are often the improved versions of the previous designs and the experience of the designer is an important factor.

Quantitative considerations Quantitative selection of the best alternative is arrived at through a process of optimization. The objective function to be optimized can be an economic parameter like payout period, internal rate of return, the total annualized cost for the plant, etc. Such economic parameters are usually used in case of large equipment, process systems/plants or projects. This requires a mathematical model describing the process in terms of the process design variables namely (1) operating conditions (temperature, pressure, flow, etc.) and (2) equipment specification parameters (capacity, number of separation stages, etc.). The mathematical model describing the design will be relationships among the design variables expressed as equations and inequations (“>” or “ 700 m2/m3 or a hydraulic diameter (Dh)
6 mm for operating in a gas stream and b > 400 m2/m3 for operating in a liquid/phase change stream. A laminar flow exchanger (mesoscale heat exchanger) has a surface area density greater than about 3000 m2/m3 (100 mm
Dh
1 mm) and micro (scale) heat exchanger has b greater than about 15,000 m2/m3 (1 mm
Dh
100 mm). Examples of compact heat exchangers are plate fin, tube fin and rotary regenerators for gas flow on one or both the fluid sides and gasketed, welded or brazed heat exchangers and printed circuit exchangers for liquid flow. The basic flow arrangement in compact heat exchangers are single-pass cross-flow, counterflow and multipass cross-counterflow. The tubular and plate-type exchangers with b less than 700 m2/m3 give a heat exchanger effectiveness around 60% or less. Effectiveness (ε) of a heat exchanger is defined as the ratio of the actual heat transfer rate to the maximum possible heat transfer rate thermodynamically permitted. This is discussed in greater detail later in this chapter. For a much higher effectiveness (around 98%), a more compact surface is required. Fins are usually added to increase surface area and exchanger compactness for the same temperature difference. Depending on the design, they can increase the surface area by 5e12 times the primary surface area and the resulting exchanger is referred to as an extended surface exchanger. Fins while increasing the heat transfer area may or may not increase the heat transfer coefficient. Interrupted fins (strips, louvers, etc.) increase area as well as heat transfer coefficient. The increase can be two- to four-fold. Usually an increase in fin density reduces the heat transfer coefficient associated with fins. Plate fin and tube fin geometries are the two most common types of extended surface exchangers. Internal fins in tube are less common. Mostly low finned tubes are used in shell and tube exchangers to increase the surface area on the shell side when the shell-side heat transfer coefficient is low. Highly viscous liquids, gases or film-wise condensing vapours on the shell side cause low heat transfer coefficient. Fins add to structural strength. Fins/studs may also be used to aid thorough mixing of a highly viscous liquid. The low finned tubes usually have helical or annular fins. Double-pipe exchangers usually employ longitudinal fins. Fins on the inside of the tube are either integral fins or attached fins. The fin efficiency increases with

24

Chapter 2 Heat transfer processes in industrial scale

channel

heat pipe

fin

Hot fluid channel Separator plate

Cold fluid channel

FIGURE 2.4 Heat pipe heat exchanger.

(i) decreasing annular heat transfer coefficient, (ii) increasing fin thermal conductivity and (iii) decreasing fin size. Finned tubes are usually unsuitable for fouling and corrosive shelleside fluid. An air-cooled exchanger is a finned tube exchanger in which the hot process fluid (liquid or condensing vapour) flows inside the tubes and atmospheric air is circulated by forced or induced draft over the outside extended surface. The airflow path is kept short through a layer of tubes and the face area is kept large to keep the fan power low. A heat pipe is a closed tube or vessel with the inner surface usually lined with a capillary wick (porous lining, screen or internally grooved wall). A heat pipe heat exchanger (Fig. 2.4) comprises of a bundle of heat pipes which are evacuated and partially filled with a heat transfer fluid Heat Pipe Heat Exchanger (working fluid sufficient to wet the entire wick). Hot and cold fluids, usually gases, flow continuously across separate parts of the exchanger. It is also possible to transfer heat from a hot to a cold solid by embedding the two ends of the heat pipe exchanger in the two solids. Heat transferred to the hot end of the heat pipe vaporises the heat transfer fluid inside the pipe. The vapour travels to the condensing end where it condenses by transferring heat. The condensed liquid returns to the evaporator section by the capillary action of the wick and/or gravity. A well-designed heat pipe will operate (transfer heat) as long as there is temperature difference between the hot and cold sections. Usually the temperature difference between the evaporating and condensing section is small (w5 C), thus reducing the overall thermal resistance. In gasegas heat exchangers the heat pipes are usually finned.

2.2.2 Regenerator Regenerative exchangers are exclusively used for gas to gas sensible heat transfer, e.g., in waste heat recovery, dry and moist air heat exchange in air driers, etc. In regenerators, both fluids flow alternately through the same passage. The heat transfer surface is a cellular structure, referred to as matrix or a porous solid bed. The matrix picks up the heat from the hot fluid and later transfers the same to the cold

2.2 Exchanger types

25

fluid when it flows through the matrix. Hence, the heat transfer is not unidirectional as in recuperators. To operate with continuous flow of streams and limit the periodic temperature variation of the fluids, either the matrix is physically moved periodically into and out of the fixed stream of gases (rotary regenerator) or the gas flow is diverted using valves to and from the fixed matrices (fixed matrix regenerator). A small amount of fluid is always trapped in the matrix that gets mixed with the other fluid stream on switching of the fluids. Also a small leakage of the higher pressure fluid to the lower pressure fluid is expected in real systems. Therefore, it cannot be used for systems where contamination of one fluid by the other is unacceptable. In air heating applications, humid air may transfer moisture up to about 5% to dry air. The advantages of regenerators over recuperators are (a) compactness e smaller exchanger for given exchanger effectiveness and pressure drop, (b) cheaper option, (c) simpler inlet and outlet header design for distribution of gases in the matrix and (d) can work even with particulate laden gases that cause fouling in recuperators.

2.2.3 Fluidised bed exchanger In a fluidised bed exchanger, usually the shell side of a two fluid exchanger contains the (fluidised) bed of fine particulates, e.g., a tube bundle immersed in a fluidised bed of sand or coal particles as schematically shown in Fig. 2.5. When the bed gets fluidised, there is a thorough mixing of the particles and a nearly uniform temperature in the bed. Much higher heat transfer coefficient is achieved on

Hot flue gases to particle removal and heat exchanger Heat transfer tubes Steam

Solid fuels feed

Limestone Water

Fluidizing air

Ash

FIGURE 2.5 Schematic representation of a fluidized bed boiler.

26

Chapter 2 Heat transfer processes in industrial scale

the fluidised side compared to particle free or dilute phase (particle laden) gas flow. In coal-fired fluidised bed combustors, the fluid bed exchanger is an efficient way of heat extraction from the bed. Typical applications of fluidised bed heat exchanger are drying, mixing, adsorption, reaction, coal combustion and waste heat recovery.

2.2.4 Direct contact heat exchanger In a direct contact heat exchanger, the hot and the cold fluids come in direct contact and exchange heat. Most direct contact heat exchangers involve mass transfer in addition to heat transfer as in evaporative cooling (cooling tower), cooling of hot gases with water spray, barometric condenser, etc. The enthalpy of phase change is usually predominant in such an exchange. Direct contact exchangers are advantageous due to (A) very high volumetric heat transfer rates, (B) inexpensive construction due to minimal hardware requirement, (C) absence of heat transfer surface between the fluids and (D) no fouling. However, these can be used only for services where direct contact between the streams is allowable. Cooling tower discussed in Chapter 7 involves such direct contact heat transfer.

2.3 Flow arrangement Each fluid in the exchanger may flow in single or multiple pass. A fluid makes one pass if it flows once through the full length of a section of the heat exchanger. If the fluid subsequently reverses its flow direction after full length flow and flows again through the same section, it makes a second pass. Multipass arrangements in shell and tube exchangers are elaborated in Chapter 4.

2.3.1 Countercurrent flow exchanger For single-pass exchangers, the flow of the fluids is usually co-current (parallel to each other in the same direction) or countercurrent (parallel to each other but in the opposite direction). A higher effective temperature difference results in case of countercurrent flow for the same inlet temperature of the fluids and heat duty of the exchanger. This leads to lower heat transfer area and a smaller exchanger as long as the overall heat transfer coefficient is nearly the same. In addition, the maximum temperature difference across the exchanger tube wall either at the hot or the cold fluid end (for an equivalent performance) is the lowest and that produces lower thermal stress in countercurrent flow as compared to the other flow arrangements. There are manufacturing difficulties associated with true counterflow arrangement in plate fin exchangers.

2.3.2 Co-current flow/parallel flow exchanger Co-current flow has the lowest exchanger effectiveness among single-pass exchangers for a given overall thermal conductance (U  A e explained later), fluid flow rates and temperatures. These are preferred when the cold fluid has high viscosity. The entering cold fluid at lower temperature exchanges heat with the entering hot stream at high temperature and gets heated quickly. This reduces its viscosity and improves the cold fluideside heat transfer coefficient. This also reduces the pressure drop for the cold fluid. However, the thermal stress in the exchanger at the inlet is higher due to the higher temperature difference between the fluids.

2.3 Flow arrangement

1

1

2

(A)

27

1

2

(B)

2

(C)

FIGURE 2.6 Cross-flow arrangements: (A) 1 and 2 both unmixed (B) 1 unmixed, 2 mixed (C) 1 and 2 both mixed.

2.3.3 Cross-flow exchanger Compact exchangers usually employ cross-flow arrangement (two fluids flowing normal to each other). Cross-flow is further classified as unmixed and mixed flow. In Fig. 2.6, (a) the cross-flow is ‘unmixed’ as the fins force the fluid to flow through a set of interfin spacing and prevent it from moving in the transverse direction parallel to the tubes. The cross-flow in (b) is ‘mixed’ as the fluid is free to move in the transverse direction as well. Both fluids are mixed in the arrangement (c). Unmixed flow occurs in a car radiator. Mixing in the fluid significantly affects the heat transfer characteristics of a heat exchanger. The thermal effectiveness for cross-flow exchangers lies between co-current and countercurrent flow arrangements. However, if the desired exchanger effectiveness (this is elaborated later in this chapter) is above 80%, the cross-flow exchangers may become uneconomic due to its large size. In such cases, the counterflow arrangement is preferred. Cross-flow pressure drop is lower due to shorter hydraulic path of the cross-flow fluid. Extended surface exchangers use cross-flow as it ensures simple header design. Plate-type exchangers and trombone coolers employ cross-flow arrangement. Crossflow exchangers are not very common in process industry.

2.3.4 Split flow exchanger This is observed in TEMA G shell of a shell and tube exchanger (Fig. 4.5) where the shell-side fluid enters at the centre of the exchanger and divides into two streams. Each stream flows along the exchanger length over a longitudinal baffle, makes 180 degrees turn at the end and then flows back longitudinally to the centre. The streams unite at the centre and leave through the central nozzle. The tube-side fluid flows straight through the tubes.

2.3.5 Divided flow exchanger In this exchanger (TEMA J shell, Fig. 4.5), the shell-side stream divides into two after entering at the centre of the shell. Each stream flows longitudinally along the exchanger length and exits from nozzles provided at each end of the exchanger. The tube-side fluid flows straight through the tubes.

2.3.6 Multipass exchanger When the heat exchanger design results in very long lengths or very low velocities or a low effectiveness, a multipass exchanger is the option. Single-pass exchangers can be arranged to form such

28

Chapter 2 Heat transfer processes in industrial scale

Shell fluid

Th,in

CONDENSING

Tube fluid

Th,in

Th,out

COOLING Th,out

Tc,out Tc,in

EVAPORATING

a) BOTH FLUIDS CHANGING PHASE

Th,in

CONDENSING

HEATING

Tc,out

Th,out

DE-SUPERHEATING SUBCOOLING CONDENSING

Tc,out HEATING

Tt

(B)

Tc,out

Tc,in

Shell fluid

HEATING

b) ONE FLUID CHANGING PHASE

Tc,in

f) ONE FLUID CHANGING PHASE

Tube fluid

Th,in

Th,in COOLING

COOLING

Tc,out

Th,out

Th,out

EVAPORATING HEATING SUPERHEATING

T EVAPORATING c,out

c) ONE FLUID CHANGING PHASE

Th,in

Tt

e) COUNTERFLOW , NO PHASE CHANGE

Th,in

Tc,in

Ts

Tc,in

ONE FLUID CHANGING PHASE

COOLING

Th,in

Th,out

Ts

PARTIAL CONDENSATION

Tc,out Th,out

Tc,out Tc,in

HEATING

d) PARALLEL FLOW , NO PHASE CHANGE

HEATING

Tc,in

Tt

CONDENSABLE AND NON-CONDENSABLE COMPONENTS

(A)

(C)

FIGURE 2.7 Temperature profile of fluids in different cases.

multipass arrangement. There are also multipass flow arrangements which have no counterpart in single-pass flow. Multipass shell and tube exchangers may exhibit parallel flow, counterflow, split flow or divided flow. Temperature profile of fluids for different configuration of exchangers with and without phase change is shown in Fig. 2.7. Multiple passes increase exchanger effectiveness at the cost of increased pressure drop.

2.4 Exchanger selection The aforementioned discussion highlights the variety of available heat exchangers and the large number of dimensional variables associated with different components of a shell and tube exchanger or surface geometries for plate-type, extended surface or regenerative exchangers. Regarding selection of the most suitable exchanger option, one must remember that there is rarely a single option that is best

2.5 Heat exchanger design methodology

29

(optimum) for a given application. Near-optimum heat exchanger designs involve several trade-offs. For example, a cheaper exchanger can be used at the cost of reduced performance and durability or a higher performance can be obtained for a heavier or more expensive exchanger. Similarly, a smaller heat exchanger can be opted at the cost of lower performance or higher pumping power for higher fluid velocities, etc. The designer needs to arrive at the optimum exchanger for a given application to meet the design requirements and constraints. Prior experience is the best guide for selection and design of an exchanger if one or more exchangers are in service for similar applications. Table 2.1 suggests some selection criteria. It needs to be understood that the values quoted are based on general industrial practices and information from different sources. These can be altered judiciously for specific design cases. To summarise, the most versatile exchanger for a broad range of operating pressure and temperatures are shell and tube exchangers for medium to high heat duties and double-pipe exchangers for lower heat duties ( 0:05; i ¼ 1; .; N; then re-estimate DTeff ;i ¼

DTeff N P

 ðqi = Ui Þ; go to step xi

ð1=Ui Þ

i¼1

xiv) Recalculate V from the following set of N equations formed by energy balance around each effect e 1st Effect        V1 ¼ F  hF  hL;1 þ S  hsteam ðTsteam ; Psteam Þ  hcon;1 = hV;1  hL;1

6.7 Design illustration

197

2nd to Nth Effect        Vi ¼ Li1  hL;i1  hL;i þ Vi1  hV;i1  hcon;i = hV;i  hL;i ; i ¼ 1; .; N

xv) Check validity of Vi estimates in step v     If max ABS Vi SNi¼1 ðVi Þ N > 0:01, i ¼ 1, .,N, then re-estimate Vi ¼ values from step (xiv), i ¼ 1, N 1, VN ¼ Vtotal 

N 1 X

Vi

i¼1

go to step v. xvi) Report e Overall steam economy ¼ (F  LN)/S. 1st Effect steam economy ¼ V1/S. iih Effect steam economy ¼ Vi/Vi1, i ¼ 2, .,N. Capacity: Feed processed per unit mass steam required ¼ F/S. Product produced per unit mass steam required ¼ LN/S.

6.7 Design illustration Design example 1 Design a forward feed triple effect evaporator to concentrate 14% w/w caustic soda solution to a product with 40% w/w NaOH. Feed liquor is available at 6 kg/s at 75
C. Last effect of the evaporator can be connected to an existing vacuum system and operated at a pressure of 7 kPa (abs). Steam with negligible superheat is available at 120
C. Estimated overall heat transfer coefficients for the effects are 3000, 2000 and 1250 W/(m2C), defined with respect to the difference in temperature of the liquid in the effect and the condensing temperature of steam/vapor heating it.

Process design deliverables i) Steam consumption (S kg/s) ii) Heat load (q W); Heat transfer area (A m2); Operating temperature (Ti
C) and pressure (Pi kg/ cm2(g)) and condensate temperature (Tcon,i
C) in each effect. iii) Flows rate of all streams (S,V,L kg/s); Concentration of NaOH in each liquor stream (x fraction w/w) iv) Enthalpy of all streams (hs, hv, hcon, hF, hL kJ/kg) Data N ¼ 3; F ¼ 6 kg/s; xF ¼ 0.14; x3 ¼ 0.4; TF ¼ 75
C; P3 ¼ 7 kPa ¼ 7/101.3 e 1.03,323 ¼ 0.964 kg/cm2 (g)

198

Chapter 6 Evaporators

BPE table for NaOH. X, fraction, w/w

0

0.1

0.2

0.3

BPE,
C

0

2.2

7.8

13.9

0.35

19.4

0.4

26.1

0.45

36.1

0.5

41.7

0.55

48.9

0.6

55.6

0.65

66.7

0.7

76.7




Tsteam ¼ 120 C; U1 ¼ 3000; U2 ¼ 2000; U3 ¼ 1250; Psteam ¼ Saturation pressure corresponding to (Tsteam ¼ 120
C) ¼ 0.9926 kg/cm2 (g) Tcon,4 ¼ Saturation temperature of steam at pressure P3 ¼ 38.4
C L3 ¼ FxF /x3 ¼ 2.1; Vtotal ¼ F  L3 ¼ 3.9 Trial 1: V1, V2, V3 are assumed equal to one third of Vtotal V1 ¼Vtotal/3 ¼ 3.9/3 ¼ 1.3; V2 ¼ 1.3; V3 ¼ 1.3; Step - A L1 ¼ F e V 1 ¼ 6 e 1.3 ¼ 4.7; L 2 ¼ L 1 e V2 ¼ 4.7 e 1.3 ¼ 3.4; x 1 ¼ FxF /L 1 ¼ 0.1787; x 2 ¼ FxF /L 2 ¼ 6(0.14/3.4) ¼ 0.2471; Estimating BPE corresponding to x 1, x 2 and x 3 in the effects from BPE table, in
C: 6.49, 10.4 and 26.1. Total BPE ¼ 42.99 Total DTav ¼ 120 e 38.4 ¼ 81.6
C
DTeff,total ¼ 81.6 e 42.99 ¼ 38.61 C Estimating individual contribution by assuming same q and Ai in all effects DTeff,1 ¼ 38.61  (1/3000)/(1/3000 þ 1/2000 þ 1/1250) ¼7.89 DTeff,2 ¼ 38.61  (1/2000)/(1/3000 þ 1/2000 þ 1/1250) ¼ 11.82 DTeff,3 ¼ 38.61 e 7.89  11.82 ¼ 18.9 T3 ¼ Tcon,4 þ BPE3 ¼ 38.4 þ 26.1 ¼ 64.5 Tcon,3 ¼ 64.5 þ 18.9 ¼ 83.4 T2 ¼ Tcon,3 þ BPE2 ¼ 83.4 þ 10.4 ¼ 93.8 Tcon,2 ¼ 93.8 þ 11.82 ¼ 105.62 T1 ¼ Tcon,2 þ BPE1 ¼ 105.62 þ 6.49 ¼ 112.11 Tcon,1 ¼ saturation temperature of steam ¼ 120
C Estimation of effect pressure (g) P1 ¼ saturation pressure corresponding to 105.62
C ¼ 0.2263 kg/cm2(g) P2 ¼ saturation pressure corresponding to 83.4
C ¼ L0.4792 kg/cm2(g)

6.7 Design illustration

199

Enthalpies hsteam(Tsteam, Psteam) ¼ 2705.9; hcon,1 ¼ hsat steam(Tcon1) ¼ 503.9 hV,1 ¼ hsteam(T1,P1) ¼ 2697.1; hcon,2 ¼ hsat steam(Tcon2) ¼ 442.9 hV,2 ¼ hsteam(T2,P2) ¼ 2669.5; hcon,3 ¼ hsat steam(Tcon3) ¼ 349.3 hV,3 ¼ hsteam(T3,P3) ¼ 2676.1; Enthalpy of solution found by neglecting heat of solution and considering NaOH (Specific heat of solid ¼ 1.4915 kJ/kg) to have enthalpy “zero” at 0
C. hF ¼ hF (75,0.14) ¼ 285.7 hL,1(112.11,0.1787) ¼ 416.2 hL,2 (93.8,0.2471) ¼ 330.5 hL,3 (83.4,0.4) ¼ 200.5 S ¼ (V1 hV,1 þ L1 hL,1  FhF)/(hsteam  hcon,1) ¼ 1.703 q1 ¼ S.(hsteam  hcon,1) ¼ 1.701(2705.9  503.9) ¼ 3735.5 q2 ¼ V1.(hV,1  hcon,2) ¼ 2931.5 q3 ¼ V2.(hV,2  hcon,3) ¼ 3016.4 A1 ¼ q1 / [U1 (Tcon,1  T1)] ¼ 0.1580 A2 ¼ q2 / [U2 (Tcon21  T2)] ¼ 0.1240 A3 ¼ q3 / [U3 (Tcon,3  T3)] ¼ 0.1276 Aavg ¼ (A1 þ A2 þ A3) / 3 ¼ 0.1365 % deviation from Aavg ¼ 5.2433, 3.0596 and 2.1837. Recalculating DTeff,1, DTeff,2 and DTeff,3 X ðqi = Ui Þ ¼ 5.124 i¼1;3

DTeff ;1 ¼ DTeff ;total  ðq1 = U1 Þ=5:124 ¼ 9.3829 DTeff ;2 ¼ DTeff ;total  ðq2 = U2 Þ=5:124 ¼ 11.0451 DTeff ;3 ¼ DTeff ;total  ðq3 = U3 Þ=5:124 ¼ 18.1835 Recalculated A values with new values of DTeff,1, DTeff,2, DTeff,3: 0.1315, 0.1329 and 0.1328. These are within 5% from Aavg. Re-estimating V values   V1 ¼ q1 þ F $ hF  L1 $ hL;1 =hsteam ðT1 ; P1 Þ ¼ 1.3000   V2 ¼ q2 þL1 $ hL;1  L2 $ hL;2 =hV;2 ¼ 1.3269   V3 ¼ q3 þ L2 $ hL;2  L3 $ hL;3 =hV;2 ¼ 1:2731 Recalculated Vi values differ from corresponding V values by more than 1%. Hence, new assumed values for are V are replaced by Vi values and calculations from Step A are repeated till the Vi and V values differ within 1% and the design calculation converges.

200

Chapter 6 Evaporators

The final converged results Parameter

Effect 1

Effect 2

Effect 3

Feed/liquor entering, kg/s

6

4.7

3.3731

Steam/heating vapor entering, kg/s

1.6901

1.3

1.3269

Solute in feed/liquor entering, (wt fraction)

0.14

0.1787

0.2490

Vapor leaving, kg/s

1.3

1.3269

1.2731

Liquor leaving, kg/s

4.7

3.3731

2.1

Solute in liquor leaving, (wt fraction)

0.1787

0.2490

0.40

BPE,


C

Effect temperature,
C Condensing temperature in steam chest,
C Condensing temperature of vapor leaving effect 3, 2

6.49

10.51

26.1

110.7

93.3

64.5

120

104.2

82.8


C

38.4

U, W/m .s

3000

2000

1250

q,W

3704.7

2936.3

3080.9

Area, m2

0.1333

0.1346

0.1345

Feed/liquor entering

285.67

410.98

328.31

Liquor leaving

410.98

328.31

200.46

Enthalpy, kJ/kg

Vapor leaving

2695.9

2668.8

2675.2

Steam/heating vapor entering

2705.9

2695.9

2668.8

Condensate leaving

503.8

437

346.8

Design example 2 Design a feed forward evaporator system to concentrate 50,000 kg/hr feed of 7% w/w glycerol available in storage tank at 27
C, to 80% w/w finished product. Steam is available at 103.6 kPa(abs) and the vacuum utility can generate 92.5 mm Hg (abs.) pressure in the last effect. The overall heat transfer coefficients in W/m2.
C1 are estimated to be 2350, 1250, 1120, if a 3-effect system is used. If more effects are used, then conservatively the first three effects be considered to have the same U and the later ones to have 1120 W/m2.
C1. Specific heat of pure glycerine is 2.42 kJ/(kg.
C).

Deliverables Designs with 3, 4 and 5 effects along with a comparison table. Please see Example of 1 for the deliverable items. Data. For 3-effect design. U1 ¼ 2350; U2 ¼ 1250; U3 ¼ 1120;

6.7 Design illustration

201

Boiling point of glycerol solutions.

Pressure (mm Hg)

Boiling point of liquor (
C)

Boiling point of water (
C)

Weight fraction of glycerol in liquor 0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

0.9564

92.3

50

50.2

50.9

52.1

53.4

55.2

57.6

61

66.2

80.1

103.1

149.19

60

60.3

61

62.2

63.5

65.5

68.1

71.5

77.3

92

117.6

233.53

70

70.4

71.2

72.4

73.7

75.6

78.5

82.2

88.3

104

132.1

355.1

80

80.5

81.4

82.6

84

86

88.8

92.8

99.3

116

146.5

525.8

90

90.6

91.5

92.8

94.2

96.3

99.3

103.5

110.3

127.8

161.1

760

100

100.7

101.6

102.9

104.5

106.7

109.6

114

121.5

139.8

175.8

The BPE table prepared based on the above table is e Boiling point elevation of liquor (
C) Pressure (mm Hg)

Weight fraction of glycerol in liquor (x)

Boiling point of water (Tw) (
C)

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

0.9564

92.3

50

0.2

0.9

2.1

3.4

5.2

7.6

11

16.2

30.1

53.1

149.19

60

0.3

1

2.2

3.5

5.5

8.1

11.5

17.3

32

57.6

233.53

70

0.4

1.2

2.4

3.7

5.6

8.5

12.2

18.3

34

62.1

355.1

80

0.5

1.4

2.6

4

6

8.8

12.8

19.3

36

66.5

525.8

90

0.6

1.5

2.8

4.2

6.3

9.3

13.5

20.3

37.8

71.1

760

100

0.7

1.6

2.9

4.5

6.7

9.6

14

21.5

39.8

75.8

The above table will be interpolated based on x and T. However we note that in last effect with 92.5 mm Hg, Boiling point of water is Tw ¼ 50
C and for x ¼ 0.8, the BPE is 16.2
C in the last effect. The first effect heating steam is at 103.6 kPa(abs) with saturation temperature of 100.6
C. If we consider three effects, the DTav ¼ 100.6 e 50 ¼ 50.6; T(3) ¼ 50 þ 16.2 ¼ 56.2
C. BPE for x ¼ 0.07 will be max 0.6
C. If we consider a driving force of w15
C in the first effect, the effect will operate at 85
C. A look at the BPE table suggests that with this range of Tw and x, the BPE estimation based only on x can entail w1
C maximum error and we decide to go for the approximation. Data N ¼ 3; F ¼ 50,000/60 ¼ 833.3 (kg/s); x F ¼ 0.07; x 3 ¼ 0.8; T F ¼ 27; P steam ¼ 103.6 kPa(abs) ¼ 103.6  .010,197 e 1.03,323 ¼ 0.02,318 kg/cm 2 (g) P3 ¼ 92.5 mm Hg(abs) ¼ (92.5/760 e 1)  1.03,323 ¼ 0.90,748 kg/cm2(g) U1 ¼ 2350; U2 ¼ 1250; U3 ¼ 1120; Tsteam ¼ saturation temperature corresponding to (0.02,318 kg/cm2(g)) ¼ 100.6
C

202

Chapter 6 Evaporators

Tcon,4 ¼ saturation temperature of steam at pressure P3 ¼ 50
C L3 ¼ FxF/x3 ¼ 72.92; Vtotal ¼ F  L3 ¼ 760.41; Trial 1: V1, V2, V3 are assumed equal to one third of Vtotal V1 ¼ Vtotal/3 ¼ 253.47; V2 ¼ 253.47; V3 ¼ 253.47; Step e A L1 ¼ FeV1 ¼ 579.86; L2 ¼ L1 e V2 ¼ 326.39; x1 ¼ FxF/L1 ¼ 0.1007; x2 ¼ FxF/L2 ¼ 0.1787 Step e B Roughly estimating BPE in effects from BPE table,
C: 0.5, 1 and 16.2. Total BPE ¼ 17.7 Total DTav ¼ 100.6 e 50 ¼ 50.6
C DTeff, total ¼ 50.6 e 17.7 ¼ 32.9
C Estimating individual contribution by assuming same q and Ai in all effects DTeff,1 ¼ 32.9  (1/2350)/(1/2350 þ 1/1250 þ 1/1120) ¼ 6.61 DTeff,2 ¼ 32.9  (1/1250)/(1/2350 þ 1/1250 þ 1/1120) ¼ 12.42 DTeff,3 ¼ 32.9 e 7.89  11.82 ¼ 13.87 T3 ¼ Tcon,4 þ BPE3 ¼ 50 þ 16.2 ¼ 66.2 Tcon,3 ¼ 66.2 þ 13.87 ¼ 80.07 T2 ¼ Tcon,3 þ BPE2 ¼ 80.07 þ 1 ¼ 81.07 Tcon,2 ¼ 81.07 þ 12.42 ¼ 93.49 T1 ¼ Tcon,2 þ BPE1 ¼ 93.49 þ 0.5 ¼ 93.99 Tcon,1 ¼ saturation temperature of steam ¼ 100.6
C % Checking if assumed BPE values match x and Tcon,2, Tcon,3, Tcon,4 values Recalculated BPE values e BPE1 ¼ BPE(93.49,0.1007) ¼ 0.6412 against 0.5 BPE2 ¼ BPE(80.07,0.1787) ¼ 1.21 against 1 BPE3 ¼ BPE(50,.80) ¼ 16.2 against 16.2 Check if for each effect, the recalculated BPE value differ from the previous value by more than 0.1
C. If the difference is more that 0.1
C, new assumed values are the recalculated values and calculations from Step B are repeated till the previous and values differ by less than 0.1
C.

Further reading 1. Badger, W. L., & Banchero, J. (1997). Introduction to chemical engineering. Bombay: Tata McGraw-Hill Edition. 2. Kern, D. Q. (1990). Process heat transfer. New York: McGraw-Hill.

CHAPTER

Industrial cooling systems

7

7.1 Introduction All process industries require streams and equipment to be cooled. Process cooling up to ambient temperature is carried out with cooling water and also with ambient air. Cooling water (CW) is the most common process utility used for picking up this heat. Once-through systems use water from a source like river, sea, or canal and after heat exchange, the warm water is returned back to the same source or some other stream. Such systems in industrial scale are less popular due to limitations like the unavailability of sufficient quantity of acceptable quality water throughout the year. Any chemical added to this water for reducing algal growth or scale reduction also leaves along with the water discharged. Circulating cooling water system is the alternative to the once-through system. Many of the industries have changed over from their older “once-through cooling water system” to circulating cooling water system and installed cooling towers. Air-cooled exchangers, also termed as fin fan coolers, are frequently used for cooling and condensing process streams with ambient air as the cooling medium. Fluid coolers are a case of circulating water surface cooler and cooling tower combined as a single unit. Drift and evaporation loss put together in a cooling tower (CT) is around 0.2% of the circulation rate. In a 10,000 m3/hr system, the same will be 20 m3/hr or 3360 m3 in a week, which is close to the size of a small pond. The operation of this tower will dry up one such Cooling Tower vs ‘Once through system’ pond full of water every week. As an alternative, if it is possible to use a once-through system, 10,000 m3/hr of water would have been returned to the environment at a temperature, which would be higher than its initial temperature by around 12
C. In the case of a flowing stream, the warm water returned mix with the rest of the water and the stream temperature, usually, rises by around 4e6
C that subsides within a reasonable distance downstream e say after flowing 500 m and becomes almost imperceptible. The foregoing analysis suggests that although the CT option is possibly more attractive from the point of economy for the industry, it may not always be so from the environmental aspect.

Process Equipment and Plant Design. https://doi.org/10.1016/B978-0-12-814885-3.00007-5 Copyright © 2020 Elsevier Inc. All rights reserved.

203

204

Chapter 7 Industrial cooling systems

The present chapter discusses the design of cooling towers (CT) in which the circulating cooling water comes in direct contact with air. The quality specification of cooling water is available in Chapter 18 (Table 18.8).

7.2 Cooling tower Wet Cooling Tower is the most common form of industrial cooling arrangement. The circulating cooling water picks up heat from the plant and rejects it in the cooling tower. Hot water entering at the top is distributed within the tower structure in a manner that exposes a large water surface to the air passing through. Water distribution is achieved by spray nozzles or distribution pans. In most cases, various types of fill media are used to increase the air-water contacting surface. Airflow can be due to natural draft or by fans forcing in or sucking out the air. As water trickles through the fill media, the air is blown across the fill to have direct contact with the falling water. A small part of the water evaporates in the tower with its heat of vaporization (latent heat) being supplied from the rest of the water that gets cooled. The cold water is circulated back to the plant. Since water comes in direct contact with air in a cooling tower, this type of cooling system is called an open recirculating cooling system. Advantages of cooling tower includes •

• • • • •

Can achieve lower water temperatures compared to other cooling methods. Achieved temperature can be lower than the ambient air dry-bulb temperature if the air is relatively dry and typically to a temperature around 4
C above the ambient air wet-bulb temperature. Relatively cheap and efficient option for rejecting low-grade heat1 from the warm cooling water returned from the process. Minimal freshwater requirement to make up water losses, thus promoting water conservation Lower environmental impact due to cut down on chemicals added to cooling water (as only a small amount of chemical is lost with the water leaving the circulating water system). Close control and improved water quality that reduce fouling and corrosion tendency in process equipment. Occupy a smaller plot space as compared to a fin fan air cooler used to remove the same heat duty.

7.2.1 Classification Cooling towers are classified based on build, heat transfer method, airflow pattern, and the principle used in creating the air draft. Fig. 7.1 shows the basis of classification and the different types of wet cooling towers.

1 Low-grade heat refers to the heat energy associated with a hot stream that has a relatively low temperature. Removing this “low-grade heat” would involve a lower temperature difference (DT) with the cooling media. This calls for a large value of “U  A” for the heat exchange process designed to remove heat Q (Q ¼ U  A  DT). In Indian condition cooling water returned from the processes is typically at a temperature of 45e47
C and the average ambient temperature is around 33
C. Therefore, the warm water is cooled by direct contact with air in cooling towers that have high “U  A.”

7.2 Cooling tower

(A)

205

Cooling Tower

Classification based on Build

Classification based on Heat Transfer method

Classification based on air draft

Classification based on air flow pattern

Natural draft

Mechanical draft

Package type

Wet

Field erected

Dry

Counterflow

Induced draft

Fluid

Crossflow

Forced draft

Atmospheric

(B)

Air Air

Drift eliminators

Fan

Drift eliminators

Nozzles Hot water

Redwood sheath and fill

Fill

Hot water

Nozzles

Fan Air

Air Louvers

Air

Basin

Cold water

INDUCED DRAFT

Basin

FORCED DRAFT

Nozzles Hot water

Louvers Air

Basin

Cold water

ATMOSPHERIC

FIGURE 7.1 (A) Cooling tower classification and (B) schematic of different types of cooling tower.

Cold water

206

Chapter 7 Industrial cooling systems

Classification by build Packaged Towers: These are common facilities for low heat rejection requirements, such as food processing plants, textile plants, hospitals, hotels, malls, chemical processing plants, automotive factories, etc. The preassembled towers are transported and fixed at an appropriate location. Due to their intensive use in domestic areas, sound level control is an important issue. Field erected towers: These are larger compared to the packaged type and common in petroleum refineries, petrochemical complexes, power plants, steel plants, fertilizer plants and other process industries.

Classification based on air draft Atmospheric Towers: These are rectangular chambers with fills having two opposite louvered walls for entry of atmospheric air driven by its own velocity. The air inside is moist and is also warmer than the ambient air due to contact with hot water. The presence of water vapor (molecular weight-18) makes moist air lighter than dry air (molecular weight-28.8) at the same temperature. This creates an updraft in the tower due to the buoyancy of moist and warmer air. As hot air moves upwards through the tower, fresh cool air is drawn in. These towers are cheap but inefficient and are not economical for high capacities. Its performance depends upon the direction and velocity of wind. Natural Draft Tower: Natural draft cooling towers are tall, up to 200 m and have a hyperbolic shape. The draft is generated by the difference in density between the ambient air and the moist, warmer air inside the tower. Atmospheric air flowing at an altitude across the top of these tall towers creates an additional draft. Fresh cool air is drawn in through the air inlet at the bottom. These towers use very large concrete chimneys to direct air through the fill media. The natural draft created is sufficient and no fan is required. There is almost no recirculation of hot air that could affect the performance in these tall towers. Natural draft towers can be crossflow type or countercurrent flow type. In crossflow configuration, the packing is external to the tower and the ambient air drawn through the fill is in crosscurrent flow to the falling water. In counterflow towers, the fill is inside and an updraft of air flows through it. Concrete is used for the tower shell supported on a set of reinforced concrete columns. The hyperbolic shape allows more packing to be fitted in the bigger area at the bottom of the shell. This shape also provides greater structural strength and also directs the air to flow smoothly towards the center, thus aiding the upward draft. The pressure drop across the tower is low and the air velocity above the packing may vary from 1 to 1.5 m/s. These cooling towers are for large heat duties and are generally used for water flow rates above 45,000 m3/hr because large concrete structures are expensive. Natural draft cooling towers are usually economical for substantially large cooling requirements in industries like steel plants. Because of their large size, construction difficulties and cost, natural draft towers have been replaced by mechanical draft towers in many installations when economics favor such an option. Mechanical Draft Towers: These employ large fans to force or draw air through the water stream. The water falls downwards over fill surfaces, which increases the contact between the water and the air. Cooling rates of mechanical draft towers depend upon various parameters, such as fan diameter and speed, fill characteristics, etc. The two different classes of mechanical draft cooling towers are e Forced draft: These have one or more fans located at the air entry to force air into the tower. The vertical fans force air at a low velocity horizontally through the packing and then vertically against the downflow of water. The drift eliminators located at the tower top disengages the water droplets entrained in the air. The fans handle mostly dry air, significantly reducing erosion and water condensation problems.

7.2 Cooling tower

207

Induced draft: The draft of hot, moist air is created by fans mounted at the air exit/tower top. This produces low entering and high exit air velocities, reducing the possibility of moist air recirculation into the air intake. This is the most common type for moderate-sized cooling towers where the liquid load is around 12e20 m3/m2 $ hr.

Classification based on airflow pattern Crossflow Towers: These have airflow perpendicular to water downflow. Air enters through one or more vertical faces of the tower to meet the fill material and flowing water (perpendicular to the air) that descends through the fill by gravity. Water distributor in these towers consists of a deep pan with holes or nozzles in the bottom. The exiting air in larger towers usually flows into an open plenum area. Crossflow towers are thermodynamically less efficient and are mostly for small capacity. These generally use induced draft, but crossflow natural draft towers are also in use. Counterflow Towers: These have airflow opposite to water flow. Air first enters an open area below the fill media and then moves up vertically through water, which flows downward by gravity through the fill. Counterflow configuration ensures higher thermodynamic efficiency and is used for large capacity towers. These can have either forced or induced draft configuration.

Classification based on the heat transfer method Wet Cooling Towers: These lower the water temperature by evaporative cooling. A small part of the water evaporates in the tower by absorbing the heat of vaporization (latent heat) from the remaining water, which gets cooled in the process. Thus cooling in a wet cooling tower is a combined heat and mass transfer process and the heat transfer is predominantly by latent heat. These are the most common types of cooling towers and shall, henceforth, be referred to as cooling towers. Dry Cooling Towers: These are basically air-cooled exchangers used for cooling circulating water by the transfer of sensible heat. Since air and water are not in direct contact, there is no evaporation loss or contamination of water, and this is particularly advantageous in areas having water shortage. The cooling water recirculates in a closed loop between the process equipment/heat exchanger and the cooler. Fluid Coolers: In these coolers, the circulating cooling water passes through a tube bundle, upon which water is sprayed and a fan-induced draft is applied to cool the external surface of the tube. This is a case of circulating water surface cooler and the cooling tower combined as a single unit. The resulting heat transfer performance is close to a wet cooling tower with the advantage of protecting the circulating cooling water from air exposure and picking up dirt. Fluid coolers avoid cooling water contamination due to exposure to atmospheric air and are used when the cooling water quality is important.2

7.2.2 Components of a typical cooling tower The schematic diagram of a typical “induced-draft countercurrent” cooling tower is shown in Fig. 7.2. The basic components are described below e Cell: Mechanical cooling towers are constructed with one or more square cells. The cells are placed side by side, sharing a common partition. Air entry/exit is through the opposite sides. The cell 2

More on the quality specification of cooling water is available in Ch. 18.

208

Chapter 7 Industrial cooling systems

FIBERGLASS FAN STACK DRIVE SHAFT FAN

GEAR REDUCER

MOTOR DRIFT ELIMINATORS

PERIMETER HANDRAIL FAN DECK HOT WATER INLET

DISTRIBUTION SYSTEM PVC FILM FILL

CORRUGATED CASING PANEL

AIR INLET COLD WATER BASIN

FIGURE 7.2 A typical induced draft counterflow cooling tower with film fill.

size is limited to 20  20 m to avoid nonuniformity of airflow throughout the cell section. Partitioning the tower in cells also helps in achieving energy-efficient tower operation at a lower load when only the required number of cells in a tower is run at their optimum rated capacity. During construction of multi-cell cooling towers, usually, a provision is kept for adding one extra cell. This takes care of increased future demand or imperfections in design that may show up as limitations in cooling. Frame and casing: Frame and casing are partitioned into cells. The structural frames support the exterior enclosures (casings), motors, fans and other components. The casing contains the water within the tower, provides an air plenum for the fan and transmits wind loads to the tower frame. Cold-water basin: The cold-water basin, located below the tower, receives the cooled water that flows down through the tower and fill. The basin, also termed as a sump, used to be underground in older installations but nowadays, it is an overground concrete tank. In some cases, there can be an additional storage sump to collect water. Thus the cold water basin has two important functions: collecting the cold water after its flow through the tower and acting as the tower’s primary foundation. Louvers: Every well-designed crossflow tower is equipped with inlet louvers whereas counterflow towers occasionally require louvers. Their purpose is to retain circulating water within the confines of the tower, as well as to equalize airflow in the fill. Louvers slant towards the inside of the tower to return any escaping water droplet into the tower. Nozzles: Nozzles are a part of the water distribution system that distributes water to wet the fill. Uniform water distribution at the top of the fill is essential to achieve proper wetting of the entire fill surface. Nozzles can be fixed and spray in a round or square pattern, or they can be part of a rotating assembly as found in some towers with a circular cross-section.

7.2 Cooling tower

209

Air inlet: It is the area of air entry into the tower. This may take up an entire side of a tower for crossflow design or be located low on the side or the bottom of the tower in counterflow design. Fans: Cooling tower fans must move large volumes of air efficiently and with minimum vibration. Generally, propeller fans are used in induced draft towers and both propeller and centrifugal fans are found in forced draft towers. Propeller fans have ability to move vast quantities of air at the relatively low static pressure encountered. They are comparatively inexpensive, may be used on tower of any size and can develop high overall efficiencies. Compared to propeller fans, the centrifugal fans can work against higher head load and are usually used in cooling towers designed for indoor installations. However, their inability to handle large volumes of air and high input horsepower requirement limits their use to relatively small applications. Airflow in propeller fans can be adjusted by varying the blade pitch and RPM in order to provide the required airflow with minimum input fan power. The optimum speed of a cooling tower fan seldom coincides with the most efficient speed of its motor and a gear reducer is often installed between the motor and the fan. Drift eliminators: Drift eliminators remove entrained water from the exiting air by forcing it to undergo sudden changes in direction. The centrifugal force separates the water drops as these hit the surface. The collected water is returned to the tower. Fill: Most towers employ fills (made of plastic or wood) to facilitate heat transfer. Fill is used to increase contact area, as well as turbulence that promotes intimate vaporeliquid contacting. This determines the efficiency of the tower. An efficiently designed fill media with appropriate water distribution, drift eliminator, fan, gearbox and motor results in lower fan power consumption as the airflow requirement is lowered. There are two types of “Fill”- Splash fill (Fig. 7.3A) and Film fill (Fig. 7.3B). In Splash fill, water falls over successively stacked “splash bars” and breaks into fine droplets, increasing the heat and mass transfer area. The surface of these bars made of wood or plastic is wet and add to the water-air contact area. These fills have low-pressure drop and are Fill types less prone to clogging but are very sensitive to inadequate support. The Film fills consist of thin, closely spaced vertical plastic surfaces, which may be flat,

FIGURE 7.3 Fill types: (A): Splash Fill, (B) Film Fill. © 2020 SPX Cooling Technologies, Inc. All Rights Reserved.

210

Chapter 7 Industrial cooling systems

corrugated, honeycombed, or may have other patterns. Water spreads over it and forms a thin film in contact with air. This provides maximum air-water interface. Film fill is more efficient since it can result in significant power savings due to lower airflow requirements. Nevertheless, its performance is affected in the case of mal-distribution of water. These are chosen for applications where the circulating water is generally free of debris that could block the fill passageways. Low-clog film fills with larger flute sizes are employed to handle water with high turbidity. This is considered a preferred option over conventional splash type fills for seawater for power savings and better performance. Originally, cooling towers were constructed primarily with wood, including the frame, casing, louvers, fill and cold-water basin. Sometimes the cold-water basin was made of concrete. Today, manufacturers use a variety of materials to construct cooling towers. Materials are chosen to enhance corrosion resistance, reduce mainteTower material nance and promote reliability and long service life. Galvanized steel, various grades of stainless steel, glass fiber and concrete, are widely used in tower construction and aluminum and plastics are adopted for some components. Frame and casing are usually made of glass fiber, while PVC, polypropylene, and other polymers are widely used for fills. When water conditions require the use of splash fill, treated wood splash fill is still used in wooden towers, although plastic splash fill is mostly used. Nozzles are usually made of PVC, ABS, polypropylene and glass-filled nylon. Aluminum, glass fiber and hot-dipped galvanized steel are commonly used fan materials. Centrifugal fans are often fabricated from galvanized steel. Propeller fans are made from galvanized steel, aluminum, or molded glass fiber reinforced plastic.

7.2.3 Cooling tower parameters The factors influencing the performance of the cooling tower are capacity, heat load, range, wet-bulb temperature, approach temperature. For a cooling tower intended to cool hot water received from the process plant at temperature Th (
C) to the cold water temperature Tc (
C), at which it is returned to the plant, the parameters of interest are described below. A diagrammatic representation showing the relationship between the various parameters is depicted in Fig. 7.4. Cooling range (R) is the difference between the inlet hot water (Th ) and outlet cold water temperature (Tc ) in ð
CÞ. Rð
CÞ ¼ ðTh  Tc Þ The range is decided not by the cooling tower but by the process it is serving. FIGURE 7.4 Diagrammatic representation of the cooling tower system and various parameters.

(7.1)

7.2 Cooling tower

211

There are two possible causes for an increase in range: - The inlet (to the cooling tower) water temperature is increased (and the cold-water temperature at the exit remains the same). In this case, it is economical to invest in removing the additional heat picked up. - The exit water temperature is decreased (and the hot water temperature at the inlet remains the same). In this case, the tower size would have to be considerably increased because the approach is also reduced (discussed later), and this is not always economical. Capacity - Heat rejected (Q kcal/hr), and water circulation rate (C m3/hr) indicates the capacity of cooling towers. Q in kcal/hr is a product of mass flow rate of water, specific heat (Cp ¼ 1000 kcal/ mt.
C for water) and temperature difference (R) in
C.
 
Q ¼ C m3=hr  Cp  R ¼ C m3=hr  Cp  ðTh  Tc Þ (7.2) Heat load imposed on a cooling tower is determined by the process being served. Typical numbers used to estimate cooling load for some common equipment/devices are presented in Table 7.1.

Table 7.1 Typical cooling capacity requirement for different equipment. Equipment

Cooling load requirement

Air compressor -

Single-stage

130 kcal/kW$hr

-

Single-stage with aftercooler

860 kcal/kW$hr

-

Two-stage with intercooler

518 kcal/kW$hr

-

Two-stage with intercooler and aftercooler

860 kcal/kW$hr

Refrigeration, compression

65 kcal/min/ton refrigeration load

Steam turbine condenser

555 kcal/kg of steam

Diesel engine, four stroke

880 kcal/kW$hr

Natural gas engine, four stroke

1525 kcal/kW$hr (for 18 kg/cm2 compression)

Wet-bulb temperature e The amount of cooling that can be obtained from a cooling tower depends on the relative humidity or the wet-bulb temperature of the air. Irrespective of the size of cooling tower, range or heat load, it is not possible to cool water below the wet-bulb temperature of air with evaporative cooling. Thus, the wet-bulb temperature of air entering the cooling tower determines the minimum operating temperature for the process or system that uses cooling water. In practice Tc is always a few degrees above Tamb;WBT , typically a minimum of 4
C. However, it is theoretically possible to cool water to temperatures below the wet bulb temperature of the ambient air since air on entering the cooling tower gets heated and can hold more moisture.

212

Chapter 7 Industrial cooling systems

It is important to note that the specified wet-bulb temperature is generally the ambient air wet-bulb temperature, but in reality, this is often affected by the discharged moist air recirculating into the tower. Hence the ambient wet-bulb temperature (the temperature in the cooling tower area) may not be same as the inlet wet-bulb temperature (the wet-bulb temperature of the air entering the tower) and the designer may adjust the wet-bulb temperature used for sizing the tower upward to account for any potential recirculation. In general, the design wet-bulb temperature selected is close to the average maximum wet-bulb temperature in summer. The design WBT (Tamb;WBT ,
C) should not be exceeded for more than 5% of the time, and it is necessary to evaluate the effects of increased wet-bulb temperatures on the tower performance. Approach to wet-bulb temperature (A) is the difference between Tc and Tamb;WBT . Að
CÞ ¼ Tc  Tamb;WBT

(7.3)

The lower the approach, the better is the cooling tower performance, but this leads to designing a more expensive cooling tower due to its increased size. An approach lower than 2.8
C is not economical, nor will it be certified by the Cooling Technology Institute (https://www.cti.org/ and see 7.2.5). Cost-effective selections are based on a criterion using an approach close to 4
C. This is not because towers are unable to achieve lower than this, but any error in measurement becomes very significant when performance is calculated at the design point. Approach temperature above 4
C results in higher cooling water temperature in process plants and reduced efficiency of heat exchange using cooling water without much savings on cooling tower. While both range and approach should be monitored, “Approach” is a better indicator of cooling tower performance. When the size of the tower has to be chosen, the approach is the most important parameter, closely followed by the flow rate. The range and wet-bulb temperature are of lesser importance. As the water gets cooled during its downflow in contact with air, the required depth of fill over which the contact takes place is governed by the range of cooling while the plan area of the fill section increases with increasing water flow rate. In simple words, the tower height is decided by the cooling range, whereas the tower cross-section is decided by the water flow capacity. Liquid/Gas ðL=GÞ ratio of a cooling tower is the ratio of the mass flow rates of inflow water ðLÞ and dry air (G). From enthalpy balance, heat removed from water must be equal to the heat absorbed by surrounding air or ðL=GÞ ¼

ðhvo  hvi Þ R

(7.4)

Where hvo and hvi are the enthalpy (kcal/kg dry air) of air-water vapor mixture at exhaust wet-bulb temperature and inlet wet-bulb temperature respectively in (
C). Effectiveness ðEf Þ is the ratio (expressed as %) between the actual range and the maximum range. Ef ð%Þ ¼ 100


ðTh  Tc Þ R ¼ RþA Th  Tamb;WBT

(7.5)

As discussed above, a cooling tower can never be 100% effective. Moreover, in summer, the ambient wet-bulb temperature is higher as compared to winter, and this limits the CT efficiency.

7.2 Cooling tower

213

Cooling towers have certain design values, but the best cooling tower effectiveness requires adjustment for seasonal variations and tuning of water and air flow rates. Adjustments can be made by water box loading changes or blade angle adjustments.

7.2.4 Cooling water circuit in a process plant A schematic diagram of a typical cooling water system with an induced draft cooling tower is shown in Fig. 7.5. Pumps supply water (at a recirculation rate C m3/hr) from the basin to the cooling water supply header that goes to the process plant. The cold water picks up heat in the heat exchangers in the process installation and joins the cooling water return header. The “warm” cooling water return header reaches the cooling tower top by its own pressure. As shown in the figure, there is only one set of pump in the flow circuit. It develops sufficient pressure to overcome the pressure drop in supply header, exchangers, return header and also the hydrostatic head to reach the cooling tower top. The cooling water supply and the return header pressure in the processing unit usually depend on the layout, i.e., system pressure drop. However, the return header pressure should not be less than 1.5 kg/cm2(g) to reach the tower top and the supply header pressure is typically around 3.5 kg/cm2(g). The major water loss from the system is the water evaporated for cooling duty (E). Theoretically, evaporation from a cooling tower is pure water, while all of the dissolved ions are left Water losses from cooling tower behind in the system. If the water loss is only due to evaporation, the concentration of the dissolved ions in the recirculating water would continue to rise until the solubility limit of each ion is reached. Further evaporative loss would start salt deposition on the heat transfer surfaces as scales. To avoid this, the concentration of critical scaling-prone ions in the circulating water is usually

FIGURE 7.5 Schematic showing a circulating cooling water system with an induced draft cooling tower.

214

Chapter 7 Industrial cooling systems

controlled by a combination of (a) bleeding off a certain portion (B m3/hr) of the recirculation water with make up of the same quantity of water having low solid concentration, and (b) adding antiscaling additives. The blowdown stream (B) leaves either the sump or the pump delivery header. The figure also shows the droplets of water entrained in the exiting air. This is called drift loss (D). There is also some water loss (L) within the process due to leakages, draining, etc. The possible sources for leakages (uncontrolled water loss from a system) are from pump seals, valves, overflow, exchanger draining, splash out, etc. Drifts and leaks are nonblowdown water losses. One may note that the loss streams take out dissolved solids and suspended solids to some extent. The makeup water stream (M) joins the sump to compensate for the water losses from the system and to Make-up water & added chemicals keep its level steady. Dissolved and suspended solids present in the makeup water stream (M) get added to the system. Solids also enter during contact of the water with air in the tower and the sump. It may also be picked up from the contacting heat transfer surfaces and piping. The rate at which (blowdown) water is bled from a system (B in m3/hr) compared with the amount of makeup water (M in m3/hr) determines the concentration ratio. This is called cycles of concentration (COC). To check the concentration ratio in a system, a soluble ion (such as silica or magnesium) that is present in sufficient quantity and is stable and easily tested is selected and monitored in the makeup water (XM) and in the recirculating water (XC). At steady state, the rates at which the solids enter are the same as the total rate of the solids leaving with blowdown stream, process plant leakage, draining and drift loss. The blowdown streamflow is increased whenever the concentration approaches the maximum tolerable limit. Chemicals for preventing corrosion, fouling, and biological growth are added directly in the cooling water sump/pump delivery header (as shown in the figure), and the pH control chemicals are added to the CT sump. Antiscaling and corrosion control chemicals are generally added in the pipeline (pump delivery header). In some installations, chlorine is dosed in the cooling water supply line daily for about an hour to control biological growth. This is called “shock dosing” of chlorine. The concentration of chlorine in the return line is closely monitored during chlorine dozing and is never allowed to exceed 1 ppm to prevent corrosion due to chorine overdose. The stream parameters to and from the cooling tower (as mentioned above and shown in Fig. 7.5) are estimated as follows. Evaporation loss (E) e Although it varies with temperature and humidity, a general rule is that for every 6
C drop in water temperature across tower, approximately 0.85% of recirculation rate (C) will be evaporated. Accordingly, the following empirical Estimation of Stream Parameters formula can be used to estimate evaporation loss: (Reference: Perry’s Chemical Engineers Hand Book)

 E m3 =hr ¼ 0:0085  ðR=6Þ  C m3 =hr (7.6)

7.2 Cooling tower

215

The evaporation loss can also be estimated from heat balance the cooling tower, i.e., the
across  amount of heat to be removed from circulating water (Q ¼ C m3 hr  Cp  R) and the amount of heat removed by evaporative cooling (Q ¼ E  l) On equating the two, we get:   C  Cp  R E ¼ (7.7) l Where, l ¼ latent heat of vaporization of water ¼ 2260 kJ/kg
Cp ¼ specific heat of water ¼ 0.238  C;
kJ/kg C is water recirculation rate in m3 hr and R is in
C

A more rigorous way of calculating E based on humidity difference of inlet and exit air and its flow rate is illustrated in the design example at the end of this chapter. Drift loss (D) is usually estimated as a percentage of recirculation (C). Splash fill towers tend to have higher drift rates than film fill towers. Drift eliminator design, unit maintenance and airflow also have an influence on the amount of drift from a cooling system. In the absence of manufacturer’s data, D may be assumed to be: (i) 0.3%e1.0% C for a natural draft cooling tower (ii) 0.1%e0.3% C for an induced draft cooling tower (iii) about 0.01% C or less, if the cooling tower has drift eliminator Makeup water (M) is estimated from a water balance around the entire system M¼E þ B þ D þ L

(7.8)

M¼E þ B þ D

(7.9)

Ideally with negligible leakage, Cycles of Concentration e The quantities of blowdown and makeup water are optimized by the Cycles of Concentration (COC). Any of the following ion concentration or even the conductivity can be used to calculate COC. (i) COC ¼ Silica in Recirculating Cooling Water/Silica in Makeup Water (ii) COC ¼ Calcium hardness in Cooling Water/Calcium hardness in Makeup water (iii) COC ¼ Conductivity of Cooling Water/Conductivity of Makeup water From a chloride balance around the system and considering that the evaporated water (E) has no salts, COC ¼ XC =XM ¼

M E ¼1þ ðB þ DÞ ðB þ DÞ

Where, XM ¼ Concentration of chlorides in makeup water (M), in ppm by weight XC ¼ Concentration of chlorides in circulating water (C), in ppm by weight

(7.10)

216

Chapter 7 Industrial cooling systems

COC usually, range from three to seven in petroleum refinery cooling towers and may be much higher in some large power plants. While a high COC reduces the makeup water requirement of the cooling tower, it allows higher dissolved solids concentration in circulating cooling water, which results in scaling and fouling of heat transfer surfaces. Holding Capacity or System Volume (HC) is the amount of water held up in the cooling water system expressed in cubic meters. This includes the holdup in the basin, additional sump if any, and all associated equipment and circulating water piping. Time per cycle is defined as the time taken for all the water held up in a system (HC) to make one trip around the recirculating loop (from the discharge side of the recirculation pump back to the suction side of the pump). Mathematically, Time per cycle ¼ HC=C

(7.11)

Holding Time Index or Half-Life index (HTI) indicates the time required to reduce the chemical or makeup water added to a system to 50% of its original concentration. It is essentially the half-life of a chemical added to the system and is estimated based on the assumption that the rate of decrease in concentration of the chemical at any instant is proportional to its concentration. The expression for calculating the holding time index usually reported in hours is e HTI ¼ 0:693

HC . ðB þ DÞ

(7.12)

The HTI is important for a chemical treatment program and is also used to determine the requirement of some biocides to achieve proper control of microorganisms.

7.2.5 Codes and standards Cooling Technology Institute (CTI), USA, established in 1950, is a body that has standardized the cooling tower design (Industrial Cooling Tower Standard - STD-203), testing (STD-202), and several other aspects. CTI codes are among the most popular codes used these days. To properly select a tower, the designer should consider towers with CTI certified listing as this is the most widely accepted standard.

7.2.6 Thermal design Quantitative treatment of cooling tower performance by separately dealing with mass and heat transfer is laborious. Therefore, the simplifying approximation of Merkel’s total heat theory has been almost universally adopted for cooling tower calculations. Briefly, Merkel’s theory states that all of the heat transfer taking place at any position in the cooling tower (dQ) is proportional to the difference between the enthalpy of saturated air at the temperature of the water (T) at that position in the tower (h0 ) and the enthalpy of the air at the same location (h). Mathematically, dQ ¼ Kaðh0  hÞdV

(7.13)

7.2 Cooling tower

217

where dQ ¼ Heat transferred by convection and evaporation for cooling of water in volume dV per unit plan area K ¼ Equivalent heat transfer coefficient. a ¼ Area of contact (m2/m3) between air and water. The contact area ‘a’ cannot be determined, and this is combined with K as Ka which refers to the unit volume of the fill. The transfer of heat dQ to air is equal to the loss of sensible heat by water. Mathematically, dQ ¼ LCp dT

(7.14)

where L ¼ water flow rate per unit area of the tower (kg/m2 hr) Cp ¼ specific heat (kJ/kg
C) dT ¼ differential change in temperature (
C) across the volume (dV, m3/m2 of plan area) Equating Eq. 7.13 and 7.14 for Cp ¼ 1 kJ/kg
C LdT ¼ Kaðh0  hÞdV The integrated form of the thermal balance equation is Z Th KaV dT ¼ 0 L Tc h  h

(7.15)

(7.16)

where V ¼ active fill volume/plan area (m3/m2) and

KaV L

¼ tower characteristic

The equation assumes L and G to be constant, but due to evaporation, this is not true in practice; however, at normal temperature levels, the error from this assumption is not significant. A standard psychrometric chart is shown in Fig. 7.6. Based on the data from the psychrometric chart, the cooling tower thermal balance plot (Fig. 7.7) is drawn with temperature as abscissa and enthalpy per unit mass of dry air as ordinate. The plot in Fig. 7.7 is based on data of the Design Illustration in Section 7.3. Line CD is the air operating line obtained from the thermal balance of heat lost by the cooling water (L) and the same picked up by air (G, kg/hr m2) in counterflow. This line has a slope of ðL=GÞ and the coordinate of point C is ðTc ; hc Þ. For a known ðL=GÞ ratio, point D can be located as ½Th ; hc þ ðL=GÞR. Assuming that the thin film of air surrounding the water droplets is always saturated, line AB is the saturated air enthalpy (h) versus air temperature plot, which is the same as the 100% RH curve on a humidity (psychometric) chart. One may note that the driving force at any cross-section through the fill is the vertical distance between the two lines (BA and CD), which is the difference between the total heat of air bulk and air film against temperature. The integral in the right-hand side of Eq. 7.16 can be numerically evaluated from the graph. In a design problem, the ambient wet-bulb temperature (Tamb;WBT ) is known and hc is the saturated air enthalpy corresponding to Tamb;WBT . Cooling of water continues as long as the operating line remains below the line of saturated air.

218

Chapter 7 Industrial cooling systems

Humidity Ratio (kg of moisture pure kg of dry air)

-10

0

10

20

30

40

50

60

0.03

0.03

0.02

0.02

0.01

0.01

0.00

0.00

-10

0

10

20 30 Dry Bulb Temperature ºC

40

50

60

FIGURE 7.6 Psychrometric chart drawn at 101.325 kPa https://www.tecquipment.com/assets/documents/downloads/ECPsychrometric-Chart-poster-A3-0617.pdf.

FIGURE 7.7 Cooling tower thermal balance.

7.2 Cooling tower

219

Cooling Technology Insitute of USA has standardized the procedure for evaluating the integral using the Chebyshev method of integration. The integral value can also be found from a nomogram (Fig. 7.8). FIGURE 7.8 Nomogram to estimate the RHS of Merkel equation.

In order to obtain the value of the integral i.e., the KaV L value for known Tamb;WBT , from any two of the parameters - Tc, Th, R, and L=G the steps are e (i) the point in the grid section corresponding to the cooling range and Tc is joined by a line with the Tamb;WBT value on the appropriate axis. The intersection of this line with the L=G axis gives the L=G value, and the corresponding KaV L is read out from the intersection of the line and the KaV scale. Lines drawn parallel to this line provide the KaV L L values corresponding to different ðL=GÞ ratios. In a design problem (refer to the problem in Section 7.3) , the L value would be known and so will be the Th , Tc and Tamb;WBT values. It is possible for different airflow rates to achieve the same cooling but in each case the KaV L will be different. For any selected G, the foregoing procedure would give us the KaV value required for the fill. Cooling tower performance is affected by the characteristics of the fill, L fill height (H), and the water (L), and the air (G) flow rates per unit tower area. These are combined empirically and published by fill manufacturers in the following form of an equation or something similar.

220

Chapter 7 Industrial cooling systems

KaV ¼ p þ q  H  ðL=GÞn L

(7.17)

Where p, q and n are constants for a specific fill. Since the equation is empirical, one needs to be careful about the units of the terms in Eq. 7.17. Table 7.2 shows the typical design range for different fill types that may be used as design guidelines. Characteristic curve for a typical industrial fill based on data of design Illustration in Section 7.3 is shown in Fig. 7.9. This relates KaV L and ðL =GÞ for 1.5 m of fill depth. The power law form of the equation plotted on log-log axes is linear. The plot also shows a curve that is derived using Figs. 7.7 and 7.8 for different ðL=GÞ values in the design problem. The intersection of the two lines is

Table 7.2 Typical values for different types of fill. Fill type

Range of L=G ratio

Effective area, a (m2/m3)

Fill height (m)

Pumping head required (m)

Splash fill

1.1e1.5

30e45

5e10

9e12

Film fill

1.5e2.0

150

1.2e1.5

5e8

Low clog film fill

1.4e1.8

80e100

1.5e1.8

6e9

FIGURE 7.9 Tower characteristic curve- determination of the design operating point.

7.3 Design illustration

221

the design operating point for the tower for the type of fill selected. The ðL=GÞ at the operating point is used to find the airflow ðGÞ to be provided by the fan. Further steps of cooling tower design are illustrated in the design illustration (Section 7.3).

7.2.7 Notes on design and operation Effect of altitude/ambient pressure: The standard psychrometric chart given in Fig. 7.6 is drawn for atmospheric pressure of 1000 mbar. When the atmospheric pressure differs from this, the chart loses accuracy. For small changes in pressure, the error is small but for appreciably lower pressures, as at high altitudes, it is necessary to apply a correction. This is because although the enthalpy of air at a particular dry-bulb temperature and absolute humidity is independent of barometric pressure, the moisture carrying ability of air is increased with reduced pressure and this alters the composition of the air/water vapor mixture at saturation. The enthalpy at saturation, therefore, increases with altitude. The effect of this increase in enthalpy improves the driving force and tends to reduce the size of the tower needed for a particular duty. However, this is counteracted by the fan, a nearly constant volume machine, delivering a lower mass flow rate due to the reduced density of air. CTI has an elaborate procedure to make pressure deviation corrections to the cooling tower performance. The details are not included in this text. Good practices: Cooling water treatment to control suspended solids and algal growth is mandatory for any cooling tower irrespective of fill media. With increasing costs of water, efforts to increase COC by cooling water treatment would help to reduce makeup water requirements significantly. In large industries and power plants, improving the COC is often considered a key area for water conservation. Drift loss is a perennial concern in cooling towers and nowadays, most of the end-user specifications assume a 0.02% drift loss. However, improved design and material (mostly PVC) being employed have improved drift eliminators with loss as low as 0.001%e0.003%. Operation of cooling tower needs to be energy efficient. Energy is spent to run the circulating cooling water system, i.e., for pumping the water and in the fans, the sum total of which should be minimized. During cold weather months, the plant engineer should maintain the design water flow rate and heat load in each cell of the cooling tower. If less water is needed due to temperature changes (i.e., the water is colder), one or more cells should be turned off to maintain the design flow in the other cells. It is a practice to run the fans at half speed or turn them off during colder months to maintain the desired temperature range.

7.3 Design illustration Design a cooling tower to cool 6000 m3/hr of warm cooling water returned at 45
C from a process plant. Cold water from the tower is circulated to the supply header at 33
C. The maximum ambient wet-bulb temperature during summer in the area where the tower is to be installed does not exceed 29
C for more than 5% of the days.

222

Chapter 7 Industrial cooling systems

Summary of available data C ¼ 6000 m3 =hr; Th ¼ 45
C; Tc ¼ 33
C; Tamb;WBT ¼ 29
C; Plant pressure drop in circulating water header ¼ 2 kg/cm2

Tower selection Water load (L) of CT usually lies between six and seven USG per ft2 (14.7e17.1 m3/m2 hr) Based on L ¼ 15 m3/m2 hr, Tower area ¼ C/L ¼ 6000/15 ¼ 400 m2 CT cells are usually square with up to 20 m arms (Refer to Section 7.2.2). In this case, four cells of dimension 10  10 m are chosen to provide the tower area. This is a moderate size tower and hence, we opt for induced draft design (Refer to Section 7.2.1). The cells are placed side by side, with air entry from opposite sides. It is also assumed that the circulating water is fairly clean, without much debris getting entrained in the flow.

Fill details The fill chosen is “Film Fill” due to its higher contact efficiency compared to “Splash Fill” and “Low Clogging Film Fill.” Typically film fill depth varies from 1.2 to 1.5 m (see Table 7.2). Among several types of film fills we choose C19 Film Fill with 1.5 m depth. Characteristics of C19 fill with depth 1.5 m under 4 to 7 US GPM/ft2 water load is given by the equation provided by the manufacturer as KaV ¼ 2:847  ðL=GÞ0:8621 L

Determination of operating L=G for the fill chosen The thermal balance diagram for the tower is based on psychometric data from a standard psychometric chart (Fig. 7.6). Data relevant to the problem, as derived from the figure, is tabulated in Table P7.1. Any intermediate value is found by interpolation. Table P7.1 Psychrometric Data for Saturated air at 1000 mbar pressure.

Aqueous Tension (Pa)

Absolute humidity (kg H2O/kg dry air)

Specific volume of dry air (m3/kg)

Specific volume of satd air (m3/ kg dry air)

Enthalpy of satd air (kcal/ kg dry air)

15

1704

0.011

0.816

0.83

10.263

20

2337

0.015

0.83

0.85

13.923

25

3167

0.020

0.844

0.872

18.206

30

4243

0.027

0.859

0.896

23.732

35

5623

0.037

0.873

0.924

31.124

40

7378

0.049

0.887

0.957

39.809

45

9585

0.065

0.901

0.995

51.005

50

12,339

0.087

0.915

1.042

66.005

Temperature (
C)

7.3 Design illustration

223

Steps of calculation The saturated air enthalpy curve (AB) is drawn in Fig. 7.7 for the range 20e50
C using data from Table P7.1. Saturated air enthalpy (on line AB) at Tamb;WBT ¼ 29
C, is read; hc ¼ 22.479 kcal/kg dry air. 1. Point C, the air entry location on the operating line is located on the graph corresponding to (Tc ¼ 33
C, hc ¼ 22.479 kcal/kg dry air) 2. A straight line CD0 is drawn that touches the curve AB. Slope of this line corresponds to the theoretically minimum air requirement, i.e., ðL=GÞmax . The slope of the operating line must be lower than ðL=GÞmax . 3. Referring to Table 7.2, we note that for “Film Fill” 1.5 < ðL=GÞ < 2 is recommended. Accordingly, a value of ðL=GÞ is assumed as say 1.5. 4. An operating line CD is drawn with a slope of ðL=GÞ. The end point D of the operating line is located corresponding to water temperature Th ¼ 45
C. Corresponding enthalpy (hD) of air exiting the tower is: hD ¼ hC þ ðL=GÞðTh Tc Þ e.g., for ðL=GÞ ¼ 1.5, hD ¼ 22:479 þ 1:5  ð45  33Þ ¼ 40:479 kcal/kg dry air R Th dT 5. The KaV L corresponding to the ðL=GÞ value assumed is the area CDD’B that represents Tc ðh0 hÞ. The integral can be evaluated by the Chebeyshev’s method as suggested by CTI. For ðL=GÞ ¼ 1.5, KaV ¼ 1.7858. L 6. For different values of ðL=GÞ, the corresponding KaV L is calculated following Step 5e7. The results are presented in Table P7.2. Table P7.2

KaV L

[

R Th

dT Tc ðh0 LhÞ,

from Merkel equation.

ðL=GÞ

1.50

1.55

1.60

1.65

1.70

1.75

1.80

1.85

1.90

1.95

2.00

KaV L

1.7858

1.8558

1.9326

2.0181

2.1138

2.2216

2.3445

2.4861

2.6515

2.8482

3.0869

ðL=GÞop corresponds to the operating condition when the chosen fill provides the required KaV L i.e., the KaV L from the Merkel equation and that from the fill characteristics match. ðL=GÞop is therefore found by drawing the lines for ðL=GÞ vs. KaV L from the fill characteristics equation and from Table P7.2 on the same graph (Fig. P7.1) and locating the intersection point. The lines intersect at ðL=GÞop ¼ 1.587 in this design example. Note: Although earlier, it has been mentioned that the tower characteristics are plotted with log-log axes, Fig. P7.1 has been drawn with the linear axis as the range of variation is small.

Fan power calculation For ðL=GÞop ¼ 1.587 and C ¼ 6000 mt/hr, L ¼ 15 m3/m2 hr, G ¼ 15/1.587 ¼ 9.45 mt/m2 hr dry airflow.

224

Chapter 7 Industrial cooling systems

FIGURE P7.1 Tower characteristics curve for the design example.

r Air flow/cell ¼ 10  10  9.4518 ¼ 945 mt/hr dry air. The induced draft fan driving the moist air needs to provide sufficient power to overcome the head losses in -

Air inlet and the rain zone (below the fill) Fill and Water Distribution Level Drift eliminator Plenum Zone

CT manufacturers have their own empirical equations to estimate each component of the losses mentioned. However, in industrial towers, the head loss in the fill and distributor level is the major component and accounts for about 70% of total head loss.

Estimating head loss in the fill and water distributor level The empirical equation for pressure drop in C 19 fill is 
DPðmmWCÞ ¼ 0:63644  exp 2:445307  103  L  ðVair Þ1:771  ðHÞ0:77  ðrair Þ where, L is water loading (m3/m2 hr) Vair is superficial air velocity through the fill (m/s) H is fill depth (m) rair is the average density of saturated air (kg/m3) in the fill

7.3 Design illustration

225

In this design: H ¼ 1.5 m; L ¼ 15 m3/m2 hr. Exit air enthalpy (hD;op ) at the operating point, hD;op ¼ hC þ ðL=GÞop  ðTh  Tc Þ ¼ 22:479 þ 1:587  ð45  33Þ ¼ 41:523 Corresponding to hD;op , the saturated air temperature interpolated from Table P7.1 is 40.865
C. Average temperature of air inside the fill ¼ (40.865 þ 29)/2 ¼ 34.933
C Interpolating data from Table P7.1, rair (saturated at 34.933
C) ¼ 1.0827 kg dry air/m3 saturated air and velocity of air Vair ¼ 1000  945=ð3600  10  10  1:0827Þ ¼ 2:4249 m=s
 DPðmmWCÞ ¼ 0:63644  exp 2:445307  103  15  ð2:4249Þ1:771  ð1:5Þ0:77  ð1.0827Þ ¼ 4:7 mmWC Fan suction flow rate ¼ 945/(4  1.0827) ¼ 218 m3/hr. Considering head loss in the fill zone to be 70% of total head loss. Head to be developed by the fan ¼ 4.7/0.7 ¼ 6.71 mmWC, Say, 7 mmWC. Interpolation of data in Table P7.1 corresponding to saturated air at exit temperature (40.865
C) gives specific volume ¼ 0.96357 m3/kg dry air. Gop (per cell) ¼ 945 mt/hr dry air ¼ (945/3600) mt/s. Power required by fan ¼ (1000  945/3600)  0.96357  (7/1000)  1000  9.8 ¼ 17351 W. i.e., (17351/746 ¼ ) 23.3 HP Assuming the efficiency of fan, gearbox and the motor to be 69%, 96% and 94%, Motor power ¼ 23.3/ (0.69  0.97  0.94) ¼ 37 HP.

Estimating make up water (M) requirement Evaporation loss (E) Total air inflow to tower (Gop) ¼ 3780 mt/hr dry air From Table P7.1, corresponding to the exit air temperature (40.865
C) and the ambient wet-bulb temperature (Tamb WB), the absolute humidity values are 0.05,141 and 0.025,365 kg H2O/kg dry air E ¼ Evaporation loss (all cells) ¼ 3780  (0.05141e0.025365) ¼ 98.467 mt/hr, i.e., 98.467 m3/hr (1.64% on 6000 m3/hr)

Drift loss (D) This is based on an empirical estimation of 0.02% of circulation rate, assuming efficient drift eliminators. (Refer 7.2.3) D ¼ 0:0002  L ¼ 1:2 m3 =hr

226

Chapter 7 Industrial cooling systems

Blowdown, including loss from the process plant (B) COC for cooling tower normally remains between 3 and 7. We assume COC ¼ 5. Accordingly,
 B ¼ E=ðCOC  1Þ  D ¼ 98:467=ð5  1Þ  1:2 ¼ 23:417 m3 =hr 0:39% on 6000m3=hr Hence, make up water, M ¼ B þ D þ E ¼ 23.417 þ 1.2 þ 98.467 ¼ 123.08 m3/hr, say 124 m3/hr, amounting to 31 m3/hr per cell (2.07% on 6000 m3/hr), assuming negligible leakage loss.

Pump calculations Selection of pump: We choose centrifugal pumps as per industrial practice for high flow rates, i.e., 2 nos.  3000 m3/hr capacity. The operating point of the centrifugal pump is chosen to have maximum efficiency. Selecting 2  3000 m3 =hr pumps will give us additional benefit of operating only 2 cells with a single pump running if need arises. Since centrifugal pump capacity can be turned down to 50%, even a single cell operation would be possible. The cooling water header pipe size is chosen to be 2000 NB, Schedule 40, having ID 18:81200 . This would entail a reasonable velocity of 2.3236 m/s, and the pressure loss estimated to be 0.0776 kg/cm2 per 100 m length. At the tower, the header climb is the total of air inlet height (taken to be half the cell width, i.e., 5 m), and fill and distributor height, i.e., 2 m. Hence, the total static head of 7 m is considered. For a static climb of 7 m, the header pressure at the base of the cooling tower, if kept at 1 kg/cm2 (g) is sufficient. Pump discharge pressure ¼ Pressure at CT base þ Pressure drop in process plant þ Pressure drop in header to and from CT (assumed 1 kg/cm2) ¼ 1 þ 2 þ 1 ¼ 4 kg/cm2(g), i.e., 40 mWC. Assuming pump and motor h to be 85% and 95%, Pump motor power ¼ (3000/3600)  40  1000  9.8/(1000  0.85  0.96) ¼ 400 kW

Cooling tower sump Base area of tower ¼ 4  10  10 ¼ 400 m2 Sump water depth ¼ 1.75 m, Sump freeboard ¼ 0.25 m (assumed) Sump holdup ¼ 400  1.75 ¼ 700 m3 This corresponds to holding time of ð60 700=6000Þ ¼ 7 min Note: In case the holding time needs to be increased, additional sump may be provided. The depth of the sump may be kept 2 m, and a space adjacent to the tower may be used in constructing the additional sump (Table P7.3).

Further reading

227

Table P7.3 Summary of the process design. Project name: cooling water system Tower

Induced draft water cooling tower

Tower model

———————

Type

concrete, rectangular, counterflow Design & Operating Conditions m3/hr

Circulating water flow, Hot(inlet) water temp.
C

Cold(outlet) water temp.
C Wet-bulb temp. Inlet


C

Ambient temp.
C

6000 45 33 29 –

Relative humidity %

– 3

Makeup water flow, m /hr

3

124 (31 m /hr per cell)

COC

5

Fan

1 per cell, Suction flow rate 218 m3/hr, DP ¼ 7 mmWC

Fan motor power

23.3 HP

Number of cells

4 in a row

Plan area

40  10 m

Overall tower dimension, L  B  H mm Water sump dimension, 2 L  B  H mm Casing material Fill Water distributor

40,000  10,000  7000 40,000  10,000  (1750 water level þ 250 freeboard) CONCRETE PVC C-19 fill 600 m3 Polypropylene nozzles

CW pump

2 nos  3000 m3/hr, discharge pressure 4 kg/cm2(g); run by 400 kW motor

CW header

2000 NB ANSI schedule no. 40

Further reading CTI Code: Tower: Acceptance Test Code for Water-Cooling Towers; Feb, 1990 revision. CTI Bulletin TPR-121: The evaluation of Cooling Tower Performance from Field Test Data.

228

Chapter 7 Industrial cooling systems

Cooling Tower Fundamentals e Marley e SPX (https://www.google.co.in/url?sa¼t&rct¼j&q¼&esrc¼s&source¼ web&cd¼1&cad¼rja&uact¼8&ved¼0ahUKEwjohPjFxZDVAhVKxLwKHZXGB2gQFggnMAA&url¼ http%3A%2F%2Fspxcooling.com%2Fpdf%2FCooling-Tower-Fundamentals.pdf&usg¼AFQjCNH skYtjKdM1MUkOqtkQZZfEgIM03Q). Cooling Tower Selection and Sizing (Engineering Design Guideline) e KLM Technology group (URL¼?). Water Purification Handbook (Chapter 31 e Open Recirculating Cooling Systems) - GE Power and Water - Water and Proicess Technologies (1995). Cooling Tower Efficiency Calculation (http://www.chemicalengineeringsite.com/cooling-tower-efficiencycalculations/76). Energy Efficiency Guide for Industry in Asia (www.energyefficiencyasia.org). Cooling Tower (https://beeindia.gov.in/sites/default/files/3Ch7.pdf).

SECTION

Mass transfer processes

III

A good scientist is a person with original ideas. A good engineer is a person who makes a design work with as few original ideas as possible. There are no prima donnas in engineering eFreeman Dyson

CHAPTER

Interphase mass transfer

8

8.1 Introduction Interphase mass transfer is the basic transport process involved in all separation processes based on concentration difference. In practice mass transfer is seldom stand-alone and is accompanied by heat and momentum transport and/or chemical reaction. For conventional separation processes, namely distillation, absorption, stripping, adsorption, extraction, and leaching mass transfer governs the overall rate and their design is based primarily on mass transfer and equilibrium considerations. On the other hand, in processes like drying, humidification, dehumidification, evaporative water cooling (in cooling towers), and membrane separation, the respective overall rates are not limited by mass transfer and their design is dealt with other transport processes (commonly heat transfer) as the controlling phenomena. Heterogeneous reaction systems which may involve mass transfer that sometimes maybe the rate-limiting step. Design of such reaction systems uses a different approach based on apparent/ overall reaction kinetic rate that includes the effect of mass transfer. In this section of the book, we focus only on the design of traditional mass transfer processes, that is, absorption, stripping, distillation, adsorption and extraction. Three steps are always involved in these processes - (i) creation of a two-phase system, (ii) mass transfer between phases and (iii) separation of the phases. Each mass transfer system in its basic configuration is built around an arrangement for contacting the phases and auxiliary systems for supply/removal of heat and handling the fluids. Distillation exploits the difference in volatility of the components to be separated, while absorption and stripping are separation processes based on the difference in solubility of the gaseous constituent(s) in the contacting liquid phase. Absorption refers to the transfer of one or more components from the gas to the liquid phase in which it is soluble. Stripping is the reverse operation, where the component transfer is from the liquid to the gas phase. In distillation, the vapor and the liquid phases of the same component(s) are contacted, and separation occurs by the transfer of the more volatile component(s) from the liquid to the vapor phase and the less volatile component(s) from the vapor to the liquid phase. Accordingly, all the components are present in both vapor and liquid phase during distillation. Extraction involves mass transfer between two immiscible liquid phases. An external stream, which is a liquid solvent, is added (to the feed) to create the immiscible phases. It is often used when the breaking of an azeotrope is difficult, or the volatilities of the components are too close, making the separation by distillation nearly impossible or uneconomical. An example is the separation of aromatic and paraffinic hydrocarbons of nearly the same molecular weight present in kerosene. Industrially, these aromatic compounds are separated by extraction using liquid sulfur dioxide as a solvent. Both adsorption and leaching involve solid-fluid (liquid/gas) mass transfer and are governed by fluid-solid equilibria. Adsorption is a surface Process Equipment and Plant Design. https://doi.org/10.1016/B978-0-12-814885-3.00008-7 Copyright © 2020 Elsevier Inc. All rights reserved.

231

232

Chapter 8 Interphase mass transfer

phenomenon, and it involves the transfer of specific component(s) from the gas/liquid phase to the surface of the solid adsorbent, where it adheres due to preferential affinity. Leaching is the preferential dissolution of one or more components from a solid mixture by contact with a liquid solvent. The appropriately chosen solvent dissolves specific components so that the desired solute can be removed. Creation of the phases: In case of distillation, the contacting vapor and liquid phases are (normally) derived from the same source, that is, the process feed by addition of heat. The other mass transfer processes use a stream different from the feed to create the additional phase. The additional phase is created by the introduction of a liquid stream (absorbent/solvent) in case of absorption, leaching, and extraction; an external vapor stream that is referred to as the inert stream in case of stripping and a solid adsorbent contacting the fluid (liquid/gas) in case of adsorption. Thus, for separation of a binary mixture, the minimum number of components is two for distillation and three for the other operations (see Table 8.1), while the number of contacting phases involved is two for all the cases. Recovery and recycle: Design of systems for processes using an external stream to create the second phase (as outlined above) will usually include an additional facility for recovery/regeneration of the said stream for reuse/recycling. This is done to improve the economy of the process. Thus, while the distillation system may be complete in itself, the other processes in order to be economic will have absorbent/solvent recovery section or adsorbent regeneration section, as the case may be.

8.2 Processes and equipment Distillation separates the feed to streams that are relatively pure. This gives the distillation process an edge over the other processes, where the substance to be separated is collected in a dilute form in the auxiliary phase(s) and may require a costly recovery step. Distillation requires energy to be supplied in the form of heat that can subsequently be easily recovered/removed from the streams. Distillation, absorption, stripping, and extraction, normally use the contacting columns. The two phases are brought into intimate contact either in a batch or a continuous flow device where the two phases exhibit cross (extraction) or countercurrent flow (absorption, stripping). Several stages of interphase mass transfer are present in the same physical column with counterflowing phases. Interphase mass transfer occurs because the contacting phases are not in thermodynamic equilibrium. The column may be fitted with internals (packings and trays) to promote contacting and also to improve separation of phases (chimney trays and demister pads) for the streams leaving the equipment. The contacting equipment can also be broadly classified based on the phase dispersed. The gas phase is dispersed as bubbles in bubble columns (sparged vessels), mechanically agitated vessels and tray towers. Liquid phase is dispersed as thin films in wetted wall columns, packed bed, and as droplets in spray towers and venturi scrubbers. Tray and packed towers are the most common industrial equipment for gas-liquid and liquid-liquid contacting. They are used for both continuous and batch processes. Stirred tanks followed by settler vessels are used for extraction and leaching and also for solid adsorbents being contacted with the liquid phase. The phase that has the higher mass transfer resistance is chosen as the dispersed phase. This is done to reduce the diffusion length in the controlling phase and also to increase its interfacial area. Bubble columns are, therefore, used when the resistance is gas-phase controlled and spray towers for larger resistance in the liquid phase. Typical features of the different mass transfer processes are shown in Table 8.1. It may, however, be kept in mind that the entries in the table are only for general information and deviations/variations are possible for specific applications.

Table 8.1 Typical features of different Mass Transfer Processes Feature

Nature of contacting equipment

Options

Single stage

Processes Distillation/ Rectification

Absorption/ Stripping

Extraction

Leaching

Adsorption

x

x

x

x

x

x

x

x

Cross flow multistage cascade Counter current flow multistage cascade

x

x

x

x

x

Established Design Approach

Equilibrium based

x

x

x

x

x

Rate based

x

x

x

Mode of operation

Batch

x

x

x

Continuous

Countercurrent with cross flow over trays

x Usually countercurrent

Usually cross flow / Countercurrent

Semi continuous


Crossflow /Cocurrent (fluidised bed) x

x

Involved phases

V -L

G-L

L-L

S-L

S-L S-G

Minimum no of components

2

3

3

3

3

x

x

x

Common Equipment

Tray column Wetted wall column

x

Bubble column Packed bed (column) Spray tower





x x

x

x

x

x

Mixer (with agitation)settler Venturi scrubber Mechanically agitated continuous contactors (Rotating Disc Contactor etc.)

x

x

x

x

x


x

Continued

Table 8.1 Typical features of different Mass Transfer Processesdcont’d Feature

Options

Processes Distillation/ Rectification

Absorption/ Stripping

Extraction

Leaching

Fluidised Bed




x

Percolating leaching equipment: Open tanks and vats, Diffusion batteries, Bucket elevator contactors, Screw conveyer contactors

x

Dispersed solid leaching: Agitated vessels e simple vessels and Pachuka tanks, Gravity thickeners, Continuous centrifuges

x

Vapour refers to gas phase below its critical temperature Used when heat of absorption is large e.g. absorption of hydrochloric acid vapour


Used for absorption of sulphur dioxide from furnace gas with slurry of limestone, lime or magnesia



Adsorption

8.2 Processes and equipment

235

Most equipment used for these continuous processes involve multistage contacting. There are two different design approaches for mass transfer equipment. Equilibrium Stage approach: The number of equilibrium contacting stages (NIdeal) required for achieving the desired separation is estimated. Additional stages are provided to account for nonequilibrium effects by considering a stage efficiency parameter h such that the actual number of stages, Nactual ¼ NIdeal/h. This ensures that the equipment is designed to achieve the desired performance. The detail of the procedure and a discussion on stage efficiency is provided in Chapter 11 on Distillation. This approach is adequate for binary mixtures and also for nearly ideal multi-component mixtures. Rate-based approach: The contacting efficiency (h) of a stage depends on the physical and transport properties of the phases and the species getting transferred. It also depends on the hydrodynamic conditions (turbulence and mixing) resulting from the contactor geometry and flow conditions (velocity, etc.) influencing the mass transfer coefficient (k) in the equipment. This variation in k is taken care of in the rate-based approach. Unlike the equilibrium stage approach, it considers nonequilibrium stages based on the physical details of the contacting stages (type, dimensions. etc.) and determines the number of actual stages (number of trays or packing height) required for the desired separation. Accordingly, the mass and energy balances around each equilibrium stage are replaced by separate balance equations for each phase around a stage. Although the same equilibrium and enthalpy relations are used, phase equilibrium is considered to exist only at the vapor-liquid interface on trays/packing and the enthalpies, concentrations, etc., are evaluated at the conditions of phases exiting. Entrainment, occlusion, chemical reaction(s), etc., can also be added to the model. Rate-based approach is superior to the equilibrium-based models particularly for nonideal systems and multicomponent mixtures. This is implemented in the RATEFRAC module of the process simulator ASPENþ and also in others like ChemSep Release 3.1 and CHEMCAD. The Design Problem: The common form in which a problem is posed states the requirement for designing a separation system for an available feed stream. The feed stream composition and quantity (in case of batch processing) or flow rate (in case of a continuous system) are known. Separation performance target for the design is specified in terms of parameters, such as (i) purity limits in terms of concentration of specific component(s) in the separated streams (ii) recovery % of the desired component(s) from feed (iii) limit of specific property values of the separated streams The client often imposes additional constraints that may include but are not limited to: (i) (ii) (iii) (iv) (v) (vi)

using a specific process or equipment (contactor) type using specific solvent/inert/adsorbent limitations on availability of hot and cold utilities maximum limit of solvent/inert/adsorbent and/or utilities per unit of feed processed specified range of feed throughputeturndown ratio limit maximum limit on investment required to build and operate the facility

236

Chapter 8 Interphase mass transfer

In most real life situations, the problem posed is usually ill defined, with incomplete information that the designer fills in through discussions with the client, his own experience, and literature survey. Experience of the designer is an invaluable component that helps to quickly conceive the initial system configuration. This step is primarily based on heuristics. Technical information gaps are filled later. Complete Design Solution: A complete design solution to the problem comprises of the specification of the system, process design of equipment with their details and the hydraulics of the complete plant. It should also contain the instrumentation and control scheme, as well as settings for any safety device (say, pressure safety valve)/arrangements for trips/interlocks. Execution of the following steps leads to the final design: • • •

• •

Configuring the overall system. This includes the contactor and the arrangement of the auxiliary equipment/facilities, such as solvent regeneration in the case of absorption and adsorption, etc. Selection of appropriate contacting equipment among the available options and performing the process design calculations that generate the information to estimate the cost. Deciding the economically optimum design choice. Minimizing only the cost of the separation equipment may increase the cost of the rest of the system, and hence, the complete system, including accessories/auxiliary system is optimized. This is illustrated in the optimum design of distillation column. Working out the mechanical details to generate the complete specifications and fabrication drawings. Detailing the plant hydraulics

8.3 Process design and detailed design of the equipment Once a contacting equipment type for a mass transfer process is chosen, its detailed design depends on the properties of the phases, their flow rates and process conditions. The detailed design of the contacting equipment can then be carried out independently. An example of this is in the design of the tower internals; the tray design is carried out in the same way irrespective of the process being absorption or stripping or distillation. However, in order to optimize the process, while carrying out the process design, some shortcut/quick estimation of the equipment parameters are done to estimate their costs. Designs dealt in this section of the book covers the process aspects of mass transfer equipment with the deliverables primarily being the height and diameter of the tower. The details of packed and tray towers for contacting the phases are dealt separately in Chapter 14 entitled Column and Column Internals. Since the process designs, in this section, pivot around phase equilibria, the same is briefly covered in Chapter 9. Subsequent chapters (Chapters 10e14) cover the process design of (i) absorption and stripping, (ii) distillation, (iii) adsorption and (iv) extraction systems. In order to deliver the design solution of mass transfer processes, one may be guided by the above chapters and Chapter 14 on Column and Column Internals. Inputs from sections V and VI of the book will be required for a complete design of the process system/plant.

8.3 Process design and detailed design of the equipment

Chapter 9: Phase Equilibria and Equilibrium staged separation. Chapter 10: Absorption and Stripping. Chapter 11: Distillation. Chapter 12: Adsorption. Chapter 13: Extraction. Chapter 14: Column Internals.

237

CHAPTER

Phase equilibria

9

9.1 Introduction Phase equilibrium data pertaining to the transferable component(s) is essential for design of any mass transfer process. Interphase mass transfer between immiscible phases occurs in the direction required to attain equilibrium, and its rate depends on the departure from equilibrium, i.e., how far away is the concentration of species i from the equilibrium concentration. Equilibrium concentration is independent of the amount/relative proportion of the two phases, and the locus of equilibrium concentration data generates the equilibrium distribution curve for each distributed component. The data can be presented either at constant temperature (isothermal data) or at constant pressure (isobaric data). Equilibrium distribution of the transferable component(s) represents the limiting composition(s) of the phases in each ideal stage of contact. The designer calculates the separation attainable in a mass transfer process consisting of one or more ideal contacting stages by estimating the equilibrium concentration(s) from phase equilibrium thermodynamics. Generally speaking, whenever a dynamic equilibrium is established between phases, the concentration of the species within individual phases (equilibrium concentration) is uniform and is fixed by the system temperature and pressure. In case of steady state, the species concentrations Equilibrium and Steady State and the thermodynamic parameters (temperature, pressure, etc.) may not essentially be the same at all locations within a phase, but at every location, these do not vary with time. Hence, the equilibrium condition encompasses a steady-state condition, but the converse is not true.

9.2 Representation of concentration The concentration of species i in a phase can have different representations. Usually, mole fraction or a quantity proportional to it, e.g., partial pressure of component in gas-liquid or gas-solid system is used to denote concentration in the gas phase. Traditionally the concentration of more volatile component in the liquid phase is x, and the same in vapor or another liquid phase richer in the component is denoted by y. In most situations, x and y denote mole fraction of the transferable species. Mole ratios X and Y may be used to represent moles of i per mole of phase free of component i, i.e., X ¼ x / (1  x) and Y ¼ y / (1  y).

Process Equipment and Plant Design. https://doi.org/10.1016/B978-0-12-814885-3.00009-9 Copyright © 2020 Elsevier Inc. All rights reserved.

239

240

Chapter 9 Phase equilibria

The same symbols (x, y, X, Y) are also employed at times to represent concentrations in terms of mass (weight fraction) instead of moles. Other units of concentration, such as normality, molarity, molality, or even a property value directly related to the concentration, e.g., absorbance or conductivity as may be appropriate, can also be used. The partial pressure of the ith component (symbol: pi) is also used to represent the concentration in the vapor phase. The symbol ci is used to represent concentration of species i in the liquid. It is expressed as mg/gm or ppmw (in liquid phase) or ppmv (in liquid phase) or something similar. The symbol ci can also be used to represent the concentration in the immiscible liquids or in the solid phase. Table 9.1 shows typical symbols used and what they conventionally represent.

Table 9.1 Representation of species concentration in phases. Symbol for concentration of component i

Conventionally represents

xi , yi

Mole fraction of species i

Xi, Yi

Mole ratio of species i Xi ¼ xi / (1  xi), Yi ¼ yi / (1  yi)

pi

Partial pressure of species i in vapor/gas phase

qi, ci or Ci

Concentration (mg/gm, ppmw or any other w/w unit), (ppmv or any other w/v unit) etc. Sometimes it may also be a property, e.g., color intensity of a dye solution, related to the concentration of component i. Usually q refers to concentration in the solid and c or C refers to concentration in the liquid e generally expressed as weight fraction

9.3 Representation of equilibrium Phase equilibrium is the condition at which each species has the same chemical potential in different phases; ideally, this is in the absence of chemical reaction. However, for practical purposes, the equilibrium representation in some systems is extended to cases where usually, a single species is transferred, and it reacts chemically with the components present in the destination phase. This is the case of the chemical equilibrium in a heterogeneous system. Examples of such cases are absorption of NH 3 in water or absorption of H2S or CO2 in aqueous alkanolamine (monoethanol amine, di-ethanol amine, etc.) solutions. In the case of adsorption on solids, the adsorbate molecules can be held on the adsorbent surface by van der Waals forces, as is the case of physisorption. In chemisorption, a stronger chemical bond (electron sharing) is formed. A further discussion on the two mechanisms of adsorption is provided in Chapter 12.

9.3.1 Graphical representation of equilibrium Several manual computational procedures are based on geometric constructions on the graphical representation of equilibrium. Typical graphical representations called equilibrium curves and their relevance are listed in Table 9.2. Few typical equilibrium curves are shown in Fig. 9.1.

Table 9.2 Graphical representation of equilibrium relationships. Representation

Special feature

Txy (Fig. 9.1A)

Constant total pressure

Tie lines join the equilibrium composition of the phases.

Binary VLE e distillation

x  y (Fig. 9.1B)

Constant total pressure

Temperature information unavailable. Points on the curve relate equilibrium compositions.

Binary VLE e distillation Gas solubilitya e absorber

X  Y (Fig. 9.1C)

Constant total pressure

Temperature information unavailable. Points on the curve relate the equilibrium compositions on solute free basis.

Gas solubilitya e absorber, stripper LLEa e extraction LSEae adsorption, leaching

Px y (Fig. 9.1D)

Constant temperature

Tie lines join the equilibrium composition of the phases.

Binary VLE e distillation

Hxy (Enthalpyconcentration) diagram (Fig. 9.1E)

Constant total pressure

Tie lines join the equilibrium composition of the phases.

Binary VLE e distillation

pi  x (Fig. 9.1F)

Constant temperature

Points on the curve relates the equilibrium compositions.

Gas solubilityae absorber, stripper VSEa e adsorption

q  c (Fig. 9.1G)

Constant temperature

Points on the curve relate the equilibrium compositions.

LSEa e adsorption, leaching

Ternary diagram (Fig. 9.1H)

Constant temperature

Tie lines join the equilibrium composition of the phases.

Ternary system. LLE e extraction LSE e leaching VSEa e adsorption Multicomponent systems at times represented.

V, vapor, L, liquid, S, solid, E, equilibrium. a for single component transfer.

Application/Use

9.3 Representation of equilibrium

Drawn at

241

Chapter 9 Phase equilibria

Min boiling azeotrope P = constant V+L V

T

V+L L

0.0

0.5 x, y

T

No azeotrope P = constant V

L 1.0 0.0

0.5 x, y

(D)

Max boiling azeotrope P = constant V L

1.0 0.0

T = Constant L

V+L

0.5 x, y

P

(A)

T

242

V+L 1.0

V 0.0

0.5 x, y

(B)

1.0

P = Constant

(E)

P = Constant V

x

=

y

y

L+V

Tie line

A

L T

V

0.0

L

1.0 x, y

1 0.8 0.6 0.4 0.2 0

V

λa V+L

Y

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1 x 250 200 150 100 50 0 0

10 20 30 40 50 60 70 80 90 100 X

FIGURE 9.1 Cont’d

Enthalpy (HV, HL)

y

(C)

0.5 x

L

HV

A

Tie line

λb HL

x, y

9.3 Representation of equilibrium

(F)

(G)

qmax

50°C

(Partial pressure of solute over liquid) pi

60°C

243

Adsorbate loading q

40°C 30°C

x (mol. fraction of solute gas in liquid)

Adsorbate conc. in solution, c

(H)

B 20

80 60

40 60 80

Plait point

40 20

Tie lines

S

A 20

40

60

80

FIGURE 9.1 Equilibrium Curves (A) Txy (B) x  y (C) X  Y along with corresponding x  y diagram (D) Pxy (E) Hxy and corresponding Txy diagram with tie lines (F) pi  x diagram (Hydrocarbon solubility in poly-alkene glycol) (G) q  c plot for adsorption (H) Ternary diagram for LLE with tie lines.

9.3.2 Mathematical representation of equilibrium Mathematical expressions relating equilibrium composition of components in different phases are derived by equating the expressions for chemical potential of a species present in different phases. The system may be binary or multi-component. In the simplest form, the phases are considered to be ideal mixtures. A more realistic representation is obtained by considering nonideal thermodynamics for the condensed (liquid and solid) and the vapor phase using activity coefficients and fugacity coefficients, respectively. In most cases of low-pressure systems, correcting only for nonideality in the liquid phase results in fairly accurate equilibrium relationships. Liquid phase nonideality often appears due to association of molecules, presence of ions, interaction between species, etc. The behavior of hydrocarbons in the gas phase may be considered “ideal” up to 10 atm pressure for most engineering calculations.

VLE: Distillation Ideal solutions obey Raoult’s law: pi ¼ xi  psat i

(9.1a)

244

Chapter 9 Phase equilibria

where pi is the equilibrium partial pressure of component i present in solution, xi is the mole fraction in the liquid phase, and psat i is the vapor pressure of the pure component at the same temperature. Raoult’s law is valid for chemically similar liquids or for components in large excess, i.e., as xi / 1 the prediction accuracy improves. Inclusion of activity coefficient (gi ) accounts for nonideality of the liquid phase and modifies the expression as: pi ¼ gi  xi  psat i The equilibrium vapor-phase mole fraction

(yi )

yi ¼ pi =P

(9.1b)

for both phases ideal is

¼ xi  psat i =P

(9.2a)

and for nonideality in the liquid phase is yi ¼ pi =P ¼ gi  xi  psat i =P

(9.2b)

where P is the total system pressure. Data on activity coefficients/equations to evaluate the same can be obtained from any standard textbook on phase equilibrium thermodynamics. Equilibrium data are also presented in the form of equilibrium constant Ki for component i. It is termed distribution coefficient and is commonly used in case of hydrocarbon systems. Ki ¼ yi =xi

(9.3)

Up to moderate pressure for dilute mixtures, some common expressions of Ki are shown in Table 9.3. For ideal solutions, Ki values can be obtained from pure component vapor pressure using Raoult’s law (Table 9.3). However, in reality, they vary with total system pressure, temperature, and composition due to nonideal behavior of the phases. Extensive charts, nomograms, and correlations are available for predicting K values for various components, particularly those associated with natural gas and oil refining industries. Table 9.3 Expressions of distribution coefficient (Ki ). Basis Raoult’s law Modified Raoult’s law

Expression  Ki ¼ psat P i  Ki ¼ gi  psat P i

Henry’s law

Ki ¼ Hi =P

Solubility

Ki ¼ psat i



xi P

Applicability Ideal solution and solute at subcritical temperature Moderately nonideal solution when activity coefficients (gi ) are known




Solutes at supercritical temperature and also for sparingly soluble solutes at subcritical temperature When solubility data in mole fraction (xi ) is available

P, total pressure; psat i , saturation pressure of pure component i; Hi , Henry’s law constant for component i in solution.

Another way to express vapor-liquid equilibrium data is by using relative volatility (ai;j ) of component i with respect to another component j. Relative volatility is related to distribution coefficient of the two components i and j as: ai;j ¼ Ki =Kj ¼

yi =xi yi =yj ¼ yj =xj xi =xj

(9.4)

Solubility: absorption and stripping Solubility data of a gaseous component in the liquid phase needs to be known for absorber and stripper design. Equilibrium solubility of components of a gas mixture over a liquid is often expressed in terms of partial pressure of the component (Table 9.3). An ideal dilute solution is described by Henry’s law: pi ¼ Hi  xi , for components in minute quantities (xi /0). Hi , the Henry’s law constant for component i depends on temperature but is relatively independent of system pressure at moderate pressure levels. xi  0:1 is considered as the upper limit for applicability of Henry’s law within engineering accuracy. The solubility of gas decreases with increasing temperature, and hence, the equilibrium (solubility) curves are steeper at higher temperatures, as is evident from Fig. 9.1F. Gas solubility increases with pressure, and it is possible to produce any gas concentration in the liquid by applying sufficient pressure as long as the liquefied form of the gas is completely soluble in the liquid. Solubility of a gas is affected by presence of other gases in the system and also by the presence of nonvolatile solute. The aforementioned relationships are applicable to nonreactive systems only and cannot be used for systems where the absorbed gas reacts with the solvent. Considering the advantages of linear interpolation/extrapolation, experimental solubility data are often presented as a reference substance plot. These are plotted with the pure solvent boiling point along the x-axis and the partial pressure of the solute gas along the y-axis; each line on the graph Linear equilibrium plots corresponds to a specific solute concentration. At times, the vapor pressure of pure liquid is also marked in lieu of its boiling point. This representation connects xi , T, and pi . Fig. 9.2 shows the reference substance plot for the ammoniaewater system. 100 80 60

Partial pressure ammonia, mmHg

40

Mole fraction H2O in liquid = 1.0

20

0.90 0.60 10 0.75

8 6 4

2 4

6

8 10

20

40

60

Vapor pressure H2O, mmHg

FIGURE 9.2 Reference substance plot for NH3eH2O system.

100

246

Chapter 9 Phase equilibria

GSE and LSE: adsorption The amount of substance adsorbed at equilibrium depends on: •

• • •

Temperature e Since adsorption is an exothermic process, higher temperature reduces adsorbate loading at constant solute partial pressure for gas adsorption. The same is also true for liquid adsorption, but the effect is much less. The temperature effect is neglected in water treatment and ambient vapor phase applications. Solute partial pressure e for gas adsorption, higher pressure increases loading at a constant temperature Specific surface area/porosity of adsorbent - Higher porosity ensures higher specific surface and larger adsorption capacity per unit adsorbent weight. Nature of solute - vapors and gases with higher molecular weight and lower critical temperature are more readily adsorbed. Chemical differences as the extent of unsaturation also influence the extent of adsorption. Permanent gases are usually adsorbed to a relatively small extent. Molecules with higher polarity are adsorbed more readily than nonpolar molecules due to which water is more readily adsorbed than hydrocarbons.

For gas adsorption, equilibria can be expressed as e fðq; p; TÞ ¼ 0; or q ¼ f ðp; TÞ where “q” represents the concentration of an adsorbed component in the solid. Adsorption equilibria are usually represented by keeping one of the aforementioned parameters constant, i.e. • • •

Adsorption isotherms: q ¼ f ðpÞ; at constant T Adsorption isobars: q ¼ f ðTÞ; at constant p Adsorption isostere: p ¼ f ðTÞ; at constant q

Isotherms are the most common form of reporting equilibrium data for adsorption and are plotted or tabulated as capacity or loading (equilibrium concentration of the adsorbed component on the solid) versus the equilibrium concentration in the fluid phase. The solid is generally referred to as adsorbent or substrate and the adsorbed component as adsorbate/solute. Loading in the solid phase is usually expressed as adsorbed mass per unit mass of (solute free) adsorbent. It can also be represented as the amount adsorbed or the number of molecules adsorbed per unit area. For commercial adsorbents used in air driers, this is usually specified as static adsorbent capacity at 10% and 60% relative humidity, often denoted as E0.1 and E0.6. At room temperature, when gas pressure does not exceed atmospheric, adsorption isotherms for most gases are linear. The nonlinearity arises due to concentration dependence of activity coefficient of adsorbate in the fluid and the solid phase. Fluid phase adsorbate concentration is expressed as partial pressure (p) or relative humidity in the vapor phase and mass (or mole) fraction per volume (mol/m3, kg/m3 or ppm, etc) for the liquid phase. Adsorption isotherms for vapor with partial pressure as ordinate can obviously extend only up to the saturation vapor pressure at the isotherm temperature. Beyond this pressure, the vapor liquefies. This characteristic is not shown for gases above their critical temperature. Adsorption isotherms of vapors often exhibit hysteresis at least over a part of an isotherm. This phenomenon is discussed in greater detail in Chapter 12. It is interesting to note that unlike solubility curves, adsorption isotherms are not always concave to the pressure axis.

9.3 Representation of equilibrium

247

The simplest adsorption isotherm expresses loading (q) as proportional to fluid phase concentration, resulting in an expression similar to Henry’s law. q ¼ b  c Isotherm equations

(9.5)

c

is the equilibrium concentration of the solute in the mixture. There are three commonly used mathematical expressions e Langmuir, BET (BrunauereEmmetteTeller), and Freundlich isotherms to describe vapor/gas adsorption equilibria as q vs p, the equilibrium partial pressure. Table 9.4 provides the details of the equations. Table 9.4 Vapor/Gas-solid adsorption isotherms. Isotherm

Corresponding equation

Assumptions/ Considerations

Langmuir isotherm

q ¼ b1  p =ð1 þ b2  p Þ b1 e Slope of the isotherm at zero coverage (Henry’s law coefficient) b2 e Constant p e equilibrium pressure of the solute

Accounts for surface coverage, i.e., for high fluid concentration, adsorbate monolayer forms on the adsorbent surface

BET equation

1 ¼ q  ½ð psat =p Þ  1   k1 p 1  þ qm  k qm  k psat

Multilayer adsorption theory based on Langmuir model

Special features

· linear in the range 0.05 < ð p = p Þ < 0.35 applicable for · not supercritical conditions used for data · seldom correlation and representation sat



qm e loading corresponding to complete monolayer adsorption k e BET equation constant psat e Saturation pressure of the solute Freundlich isotherm

Timken isotherm

q ¼ KF  ð p ÞKc KF, Kc - constants for each solute-adsorbent pair at a fixed temperature KF depends on nature of adsorbent and adsorbate

q ¼ b1 ln KT þ b1 ln c , KT ,b1 e constants

Empirical fitting of isotherm data to a linear equation in log-log coordinates

commonly used · most isotherm, although the

· ·

equation is thermodynamically inconsistent works well for heterogeneous surfaces limited application range as it does not describe isotherm over a wide range of pressure

Chemical bonding with adsorbate

While Langmuir isotherm is theoretically justified, Freundlich isotherm is of a purely empirical character since it does not have a finite Henry’s law constant. As is evident from equation, Freundlich isotherms are linear for Kc ¼ 1, concave upward for Kc > 1 and concave downward for Kc < 1. Generally 2 < Kc > 10 represents good and 1 < Kc > 2 represents moderately difficult adsorption characteristics. Kc < 1 depicts poor adsorption characteristics and requires impractically large adsorption dosage for appreciable solute removal.

248

Chapter 9 Phase equilibria

When one component of a gaseous mixture is appreciably adsorbed over others, the adsorption isotherm for the pure adsorbate is applicable, with the equilibrium pressure being the partial pressure of the said vapor. In the case of comparable extent of adsorption of both components from a binary gaseous mixture, the equilibrium data are represented as triangular plots, similar to those used in Multicomponent adsorption liquid-liquid extraction and elaborated later in this chapter. Unlike liquid solubility, adsorption is strongly influenced by both temperature and pressure and the equilibrium diagram in these cases are typically plotted under isothermal-isobaric conditions. The reference substance method of plotting for gas-liquid solubility is also applicable to adsorption data where the adsorbate is the reference substance, provided the gas temperature is less than the critical temperature. Adsorption isostere is the relation of equilibrium concentration of adsorbate in the fluid with temperature at constant adsorbent loading. Partial presLinear plots sure, dew point, or some other form of concentration is plotted against temperature or inverse absolute temperature at specific extents of loading. The abscissa of inverse absolute temperature makes the plots near-linear, and this improves the accuracy of interpolation. A typical example is shown in Fig. 9.3.

Equilibrium partial pressure of acetone, mm Hg

400

200 kg acetone 0.30 adsorbed per kg carbon 100

0.25 0.20 0.15

80

0.10 0.05

60 40

20

30 10 200

40

50 60

400 600

1000

80 2000

100

Temperature, °C 120 140 160 180 200 4000 6000

10,000

20,000

Vapor pressure of acetone, mmHg

FIGURE 9.3 Acetone adsorption on activated carbon with loading marked on the isostere lines.

9.3 Representation of equilibrium

249

Loading (mg/g/)on activated carbon adsorbate -->

Single component liquid adsorption refers to the adsorption of a single adsorbate (solute) from a solution of inert solvent(s) in which the activity of the solvent(s) is constant. While contacting fresh adsorbent (solid) with liquid, there is an uptake of adsorbate, as well as occlusion of liquid into the pores of the solid. This occlusion also leads to an apparent level of adsorption and must be carefully Adsorption from a liquid considered by the designer as this reduces the volume of liquid recovered after contacting with solid as compared to the volume of the original contacting liquid in batch processes. The apparent adsorption depends upon the concentration of solute, temperature, nature of the solvent and adsorbent. The extent of adsorption practically always decreases at increased temperature and increased solubility in the solvent. For dilute solution, the adsorption isotherm is plotted as equilibrium solute concentration in liquid versus net solute apparently adsorbed per unit weight of adsorbent (Fig. 9.4).

102

101

100 10–2

10–1

100

101

Concentration (mg/L) in liquid phase -->

FIGURE 9.4 Isotherm of phenol from aqueous solution on activated carbon at ambient temperature.

The customary procedure to determine the apparent weight of solute adsorbed is to treat a known volume of solution (v) with a known weight of adsorbent (W). As a result of preferential adsorption of solute, the solute concentration in liquid falls from cIinitial to final equilibrium value c , both expressed as mass solute/volume of liquid. The apparent adsorption of solute is calculated as fðv =WÞ ðcInitial c Þg mass solute adsorbed per unit mass of adsorbent. This is a satisfactory measure of true loading in case of dilute solutions when the fraction of original solvent adsorbed (occluded) is small. With the apparent adsorption of solute determined over the entire range of concentrations from nearly pure solvent to nearly pure solute, curves, as shown in Fig. 9.5 result. One may

250

Chapter 9 Phase equilibria

1.0

E a

Solute concentration in solution

b

D

0.0

c – 0 + Apparent weight of solute adsorbed / weight of adsorbent

FIGURE 9.5 Equilibrium curves when both solute and solvent are adsorbed on the adsorbent.

note that for pure solvent and pure solute contacting the adsorbent, the apparent loading would be “0.” Curve “a” is the apparent adsorption isotherm in the case of a dilute solution e with negligible adsorption of the solvent on the adsorbent solid. Curve “b” results when there is appreciable apparent adsorption (removal) of the solvent from the solution. In the range D to E, the solvent is more strongly adsorbed that results in the solution getting more concentrated than the original, i.e., c > cInitial and the apparent loading becoming negative. The isotherm expressions for the gas phase can also be extended for the liquid phase adsorption with the partial pressure being replaced by a suitable measure of concentration and the units of the constants modified accordingly. Thus, over a small concentration range particularly for dilute solutions, the isotherms are frequently described by Freundlich empirical equation c ¼ K1  ½v  ðcInitial  c ÞK2 .

(9.6)

Where “v” is the volume of liquid treated per unit mass of adsorbent, cInitial and c are the initial and final concentration of solute in the liquid, and thus, the product within the bracket represents “apparent loading.” While K1 is influenced by the concentration units, the value of K2 is unaffected for dilute solutions.

LLE: extraction Unlike gases, which are miscible in all proportions, liquid solutions (binary or higher) often display partial immiscibility at least over a certain range of temperature and composition. Also, there is negligible effect of pressure on liquid-liquid equilibria provided a sufficiently high pressure is

9.3 Representation of equilibrium

251

maintained to ensure that only the liquid phase is involved. Liquideliquid extraction involves systems composed of at least three substances (components) and two phases and although the insoluble phases are chemically very different, all three components generally appear to some extent in both the phases. In such ternary systems, the equilibria are depicted by triangular diagrams, in their simple form with equilateral triangular coordinates. The advantage of using a ternary plot for depicting composition is that the three variables can be conveniently plotted in a two-dimensional graph. In the triangle, the apices denote pure components and the sides denote compositions of binary mixtures. This is illustrated in Fig. 9.6A for components A, B and S where the original binary mixture contains A and B and the partially miscible solvent “S” is added to preferentially extract B. Apices A, B and S represent pure (100% molar or mass composition) A, B and S respectively and the scales on the three sides BS, SA and AB are the respective percentages of the binary solutions, e.g., B in a solution of B and S, S in a solution of S and A, and B in liquid solution of B and A. On addition of S to a mixture of A and B, the overall composition shifts from side AB to a point M inside the triangle such that the sum of the perpendiculars from the point to the three opposite sides denote the respective % of the components in the mixture, and their sum is 100%. The concentration of each species decreases linearly with distance along the perpendicular line drawn from M to the opposite side of the triangle. An important property of the diagram is that if lines are drawn from an apex through the point M and meeting the opposite side, the marking on the axis on the opposite side is the composition of the two components in the mixture.

P

A

R

E M

0.8

O

Q 0.8

0.6

0.4

0.2

0.4

P

R

0.6 M

E

0.2 S

E,R

S 0.0

0.2

0.4 0.6 xS , yS

0.8

1.0

P

0.4 Solute fraction in raffinate

0.2

0.0

x

Solute fraction in extract

0.8

0.6

0.6

1.0

0.8

0.4

0.4 0.2

Extract curve

xB , yB

Raffinate curve 0.6

0.2

=

1.0 0.8

(C)

B

y

(B)

B

yB

(A)

0.0 0.0

0.2

0.4 0.6 xB

0.8

1.0

FIGURE 9.6 Ternary equilibrium plot: (A) Equilateral triangular plot (B) Rectangular plot (C) Distribution curve.

In relation to extraction, B is the solute that, along with “feed solvent” A, constitutes the feed phase. The extraction solvent S, often referred to as “solvent,” is partially miscible with the feed solution, i.e., the addition of a suitable amount of S to the feed generates two distinct phases that are in equilibrium. In the triangular diagram for the system (Fig. 9.6A), the zone of partial miscibility is the area of the dome-shaped region OPQ. Solute B is miscible in all proportions with both A and S. The dome, thus, bound by the two equilibrium curves, which are the solubility curves with high and low concentration of the solvent S. Any point within the triangle ABS represents a

252

Chapter 9 Phase equilibria

ternary composition and when the point is outside the dome OPQ, it represents composition of a homogeneous mixture phase. A mixture with a composition corresponding to the point M within the partial miscibility region splits into two phases that are in equilibrium. Composition of the two phases in equilibrium are E and R that lie on the equilibrium curves. The line RE is a typical “tie line” that passes through the point M with the endpoints corresponding to the equilibrium compositions in the two liquid phases. The phase composition represented by E has a higher proportion of solvent S and is termed extract phase, and R represents the “raffinate phase” composition. Tie lines are rarely parallel and gradually change their slope in one direction. As evident from Fig. 9.6A, the tie lines become shorter in length as one approaches the top part of the dome until it reduces to a point at “P.” This point of inflection near the top of the two-phase envelope is termed the plait point. It signifies the condition at which the compositions of the two liquid phases in equilibrium become identical and transform into a single phase. Thus, the “plait point” composition is obviously the limit of composition where no phase separation takes place. The plait point is ordinarily not at the maximum value of B on the solubility curve. Henceforth, the equilibrium solubility curve of the extract and raffinate phase will be referred to as the “extract curve” and the “raffinate curve,” respectively. In this book, the composition in the raffinate phase is denoted by x and that in the extract phase by y, with subscripts denoting the components. The composition can be expressed in terms of mass, mole, or any other quantity basis of each substance. The curves are different for different physical quantity selected, but the material balance, equilibrium principles, and the results obtained are the same in all cases. One may note that the composition M actually results from mixing the feed and the solvent streams, and it splits into two phases with the compositions represented by E and R. On removing S from the ternary mixture M, the binary solution regains its original composition, which lies on line AB. It is often more useful to plot LLE data in rectangular coordinates with the concentration of solute (B) plotted against the concentration of solvent (S), and the concentration of feed solvent (A) is the remaining fraction, such that all weight fractions sum to unity. Similar to triangular plots, the twophase region and single-phase homogeneous liquid region are demarcated by the extract and raffinate curves, as shown in Fig. 9.6B. Tie lines connecting the two equilibrium phase compositions are also shown. The distribution of the solute in the two phases can also be conveniently shown as a distribution curve (Fig. 9.6C) of solute B In this case, the concentration of B in the extract phases is plotted on the y-axis and that in the raffinate on the x-axis. The equilibrium curve lying above the diagonal (y ¼ x) denotes that the distribution of B favors the S-rich phase (% of B in E is more than that in R). The slope of the curve is quantitatively expressed as the distribution coefficient KDi , for component i. KDi ¼ yi =xi

(9.7)

Similar to vapor-liquid systems, KDi is found from the thermodynamic considerations by equating the chemical potential of each component in the two phases. Often distribution curves are plotted with concentration of solute in raffinate and extract phase expressed on solvent-free basis denoted as XB ¼ xB =ð1  xS Þ and YB ¼ yB =ð1  yS Þ respectively.

Further reading

253

There are two main classes of liquid-liquid equilibrium in extraction: (i) One immiscible pair of species that produces the familiar envelope shown in Fig. 9.6 and is discussed above. In the system shown, species AB and SB are completely miscible while AS is partially miscible, and the mutual saturation limits are denoted by the envelope ends (O,Q) on line AS. The more insoluble the liquids A and S, the closer will be the points to the triangular apices. This nature of LLE is more common and is preferred for extraction with solvent S. Solvents are often selected to get this. Nevertheless, the type can change with temperature, and this needs to be considered. (ii) Two pairs of immiscible species exist, and so the two-phase envelope crosses the triangular diagram like a bridge. The ternary diagram drawn at 25 C in Fig. 9.7 shows this nature e AB is completely miscible in all proportions, but there is only partial miscibility between AS and BeS. BeS becomes completely miscible at 3.5 C. At 50 C, the miscibility improves further, and we note that the two-phase envelope shrinks to a single immiscible pair.

B

A

25°C

B

S

A

35°C

B

S

A

50°C

S

FIGURE 9.7 Effect of temperature of equilibrium solubility on the ternary system.

Further reading Treybal, R. E. (1980). Mass-transfer operations. McGraw-Hill Classic Textbook Reissue Series. Reid, R. C., Prausnitz, J. M., & Poling, B. E. (1987). The properties of gases and liquids. New York: McGraw-Hill. Prausnitz, J. M., Lichtenthaler, R. N., & de Azevedo, E. G. (1998). Molecular thermodynamics of fluid-phase equilibria. Pearson Education.

CHAPTER

Absorption and stripping

10

10.1 Introduction Absorption and stripping involve gas and liquid streams which transfer components from gas feed to a liquid stream in case of absorbers and from liquid feed to a gas stream in strippers. Such Absorption and Stripping contacting in a flow system can be countercurrent or cocurrent. Countercurrent contacting has the well-discussed advantage of utilizing a higher average driving force for mass transfer that results in more compact equipment in most cases of absorption and stripping. Cocurrent flow design is rarely useful in absence of reactions or when only about one stage of contacting suffices due to high solubility of the component being transferred. Nevertheless unlike countercurrent flow, the capacity of the cocurrent contactor is not limited by flooding. This allows processing at high throughput as in venturi scrubbers. Cocurrent downward gasliquid flow through packed bed is used for catalytic chemical reactions in trickle bed reactors. It is also used for situations where (i) a rapid, irreversible chemical reaction accompanies the mass transfer process, e.g., absorption of hydrogen sulfide into aqueous sodium hydroxide, (ii) an exceptionally tall tower is built in two sections, with the second section operated in cocurrent flow to save on the large diameter gas pipe connecting the two. This chapter deals with the design of continuous absorbers and strippers where the desired function is achieved through countercurrent contacting of gas and liquid in staged contactors (tray columns) and continuous contactors (packed towers). Industrially, the more economical option between a tray column and a packed column is chosen for the specific design task. The maximum attainable separation is governed by the equilibrium concentration between the phases that are affected by the process conditionsdprimarily the operating temperature and pressure, the number of stages of contacting, and the liquid to gas flow rate ratio. Inputs to the process design in case of absorber/stripper are: • • •

Inlet flow rates of gas and liquid Composition of the inlet gas and liquid in case of absorption and stripping respectively Minimum target of recovery of the component(s)

Process Equipment and Plant Design. https://doi.org/10.1016/B978-0-12-814885-3.00010-5 Copyright © 2020 Elsevier Inc. All rights reserved.

255

256

• •

Chapter 10 Absorption and stripping

Operating pressure and maximum permissible pressure drop across the column The chosen solvent (in case of absorber)/inert gas (in case of stripper)

The last two are often decided based on economic considerations, and are, therefore, left to the choice of the designer. The design decisions are usually based on the following considerations: 1) In the case of absorbers, the solvent is selected, such that, it has a high solubility for the solute. This reduces the amount of solvent. In addition, it should be relatively cheap, stable, nonvolatile, noncorrosive, nontoxic, nonviscous, nonfoaming, and preferably nonflammable. Since the exit gas is saturated with solvent, a part of the solvent is lost. A low-cost solvent is often preferred over a costly one with higher solubility or lower volatility. The availability of suitable material of construction also influences the choice to a large extent. Recovery of the solvent, usually by distillation and at times by chemical means, is almost always necessary and needs to be included in the design and cost analysis. The more effective the regeneration process, the less costly the absorber, as the regenerated solvent is left with a lower concentration of residual dissolved solute. Water is the most commonly used solvent, oils are used for hydrocarbons, and special chemical solvents for acid gases like H2S, CO2, and SO2. There is always a loss of solvent with the exiting streams, and the designer needs to specify the equivalent make up rate. In case the solvent has several components, the loss of lighter components from the solvent is more, and therefore, the makeup stream needs to be richer in lighter component(s). A light, lean oil solvent for absorbing light hydrocarbons will have higher loss but has the advantage of high absorption capacity (moles/m3 circulated) compared to heavier lean oil. The selection of a solvent close in carbon number to the components absorbed has a higher absorption capacity, but the closeness of volatility makes the separation difficult. Therefore, optimization of the absorber needs to be done by considering the absorber and the regenerator together. As an example, lean oil with three carbon numbers heavier than the lightest component is usually common in petrochemicals/refineries, whereas a lean oil heavier by about 10e14 carbon numbers is adopted in natural gas processing units. The considerations for deciding on the stripping gas (inert) are also similar and largely dictated by the availability and economics. Steam and nitrogen are common industrial stripping agents, and as in the case of solvent selection, the overall cost of the stripper operation and the cost of recovery need to be considered for arriving at the optimum choice. 2) Height of the tower and its components, e.g., the depth of packing or the number of trays 3) The optimized liquid flow rate through the absorber and regenerator 4) Inlet and outlet temperatures of the stream(s) and amount of heat required to be removed to account for heat of solution, etc. 5) Operating pressure of the column 6) Mechanical design

10.2 Tray column

257

10.2 Tray column Fig. 10.1 shows the schematic diagram of a countercurrent absorber column with N trays and a single component being transferred from the gas to the liquid phase. The mole fraction of the component in the liquid and vapor phases are x and y, respectively, and the total molar flow rates are L for the liquid phase and G for the gas phase, as marked in the figure. L0 and G0 denote the corresponding flow rates of the nonabsorbable (solute free) components, and these remain unchanged along the column. The subscripts correspond to the tray number which the stream exits. Trays are numbered from top to bottom. Accordingly, subscripts Nþ1 and 0 are used for the external gas feed and the liquid entering the column. G1 y1

L0 x0

1

2

n

Gn+1 yn+1

Ln xn

n+1

N-1

N

GN+1, yN+1

LN, xN

FIGURE 10.1 Schematic diagram of an absorber column with N trays.

258

Chapter 10 Absorption and stripping

10.2.1 Graphical determination of the number of contacting stages Assuming that no chemical reaction occurs and a single component is absorbed from the gas to the liquid phase, the number of ideal stages is determined from a graphical construction using the operating and equilibrium curves. The operating curve relates the mole fraction of the transferable component in the outgoing vapor from a tray (say nth tray) with the composition of the liquid falling on the same tray and is arrived at by species balance across the dashed control volume marked in Fig. 10.1. It extends from the point representing the Operating curve streams entering the column to that representing the exiting streams from the nth tray and thus passes through the point having coordinates (xn, ynþ1). The linearity of the operating curve depends on the units in which the concentrations of the material balance are expressed and the graph axes. When expressed in terms of mole fractions (moles solute/mole solution) or in any concentration ratio proportional to mole fraction (partial pressure, say), the operating line is Linear operating curve nonlinear. A straight line is obtained (i) for very dilute solutions (ii) when the total quantity of each phase remains constant while composition changes owing to the diffusion of several components (binary distillation) - This forms the basis of the McCabe-Thiele method of estimating the number of trays. (iii) when the concentration is expressed as the mole ratio (moles solute/mole solute free solution) as the solute free flow rates remain unchanged throughout the column. This facilitates interpolation and extrapolation. Referring to Fig. 10.1, the equation of the operating line in terms of L0 and G0 (mol/hr flow rates of nonabsorbable components) is obtained from the species balance across the dashed control volume as x0 ynþ1 xn y1 L0 þ G0 ¼ L0 þ G0 1  x0 1  ynþ1 1  xn 1  y1 and in the generalized form L0

x0 y x y1 ¼ L0 þ G0 þ G0 1y 1x 1  x0 1  y1

(10.1)

where x and y are the compositions in the respective liquid and vapor phases leaving and entering a tray. In terms of mole ratio denoted as X and Y for composition in the liquid and gas phase, respectively, Eq. 10.1 reduces to L0 X0 þ G0 Y ¼ L0 X þ G0 Y1

(10.1a)

Note: Eq. 10.1 is nonlinear on the mole fraction plot (x-y plot), i.e., the operating line is a curve plotted on x-y axes while it is linear in the XeY plot. In an absorber design problem, the feed gas flow rate (G, G0 ) and concentration (yNþ1) are known, and so is the exit gas composition (y1). The component concentration in the absorbent liquid (xo) is also knowndit is zero for fresh liquid and can have a small value in the case of regenerated solvent. Hence, for a known value of L [L0 ¼ L(1  x)], the operating line (Eq. 10.1) can be drawn on the graph. On the

10.2 Tray column

259

same graph, the equilibrium solubility data for the solute gas in the solvent liquid can also be plotted in terms of the same concentration units. Each point on the equilibrium curve represents the gas concentration in equilibrium with the liquid at its local concentration and temperature. The position of the operating line with respect to the equilibrium curve depends on the choice of abscissa and ordinate visa`-vis the direction of mass transfer. The line is above the equilibrium curve if the mass transfer occurs from gas to liquid phase (absorption), and the composition in the gas phase is represented in the y-axis. In the case of a stripper, the operating curve is below the equilibrium curve for the same choice of axes. The number of ideal contacting stages for a design problem can be determined graphically from the operating and equilibrium curves, as discussed below. x-y plot

0.35

FIGURE 10.2 Graphical determination of the number of ideal trays.

0.3

y (mole fraction)

0.25 Operating Line with L = 1.5 Lmin

0.2

(0.0035,0.2) (0.00503,0.2) 4

Operating Line with Lmin

0.15

3

0.1

(0,0.02)

0.05

2 1

0 0

(x2,y2) 0.0035

(x1,y1) 1

2

3

4

x (mole fraction)

5

6

7 ×10–3

A typical example is shown in Fig. 10.2 for a SO2 absorber where freshwater (xo ¼ 0) is used as an absorbent to reduce SO2 concentration (mixed with air) from 20 mol% SO2 to 2 mol%. The operating line generated using Eq. 10.1 corresponds to molar flux (molar flow rate/column cross-section) on a solute free basis L0 ð¼ L0 =At Þ and G0 ð¼ G0 =At Þ of 333 and 5.18 kg mol/m2.hr, respectively, where At denotes the cross-sectional area of the column. The horizontal line drawn in the figure through (0, 0.02) corresponding to the top tray (Tray no. 1) intersects the equilibrium line at “1.” The abscissa of this point denotes the equilibrium composition (x1) of the liquid leaving tray 1, and the vertical line through “1” intersects the operating line whose ordinate corresponds to the gas composition (y2). This process of drawing the horizontal and vertical lines are continued to obtain a sequence of near triangular steps till the next horizontal line has to be drawn above yNþ1. The two corner points at the base of the lowest triangular step are, therefore, (xo, y1) and (x1,y1). The corresponding lower corner points in the next upper triangular step are (x1,y2) and (x2,y2), and so on. Corresponding to the known gas and liquid flow rates (G, G0 and L, L0 ), the figure, therefore, depicts the number of theoretical trays required to enrich

260

Chapter 10 Absorption and stripping

the gas concentration from yNþ1 to y1 or more. From Fig. 10.2, this absorber requires 2.4 ideal trays. If the overall tray efficiency of 40% is considered, the actual number of trays required is 2.4/0.4 ¼ 6 trays. It is important to note that the enrichment per tray is lower at the dilute end, as separation is more difficult. Therefore, the standard practice of graphical construction that produces a more accurate estimation of the number of trays is to start from the dilute end and continue up to the final concentration.

Minimum required liquid flow rate (Lmin) in case of absorber for a given gas rate (G,G 0 ) The foregoing geometrical construction to determine the number of ideal stages of contacting could be done, only because (1) the point on the operating line corresponding to yNþ1 is situated above the equilibrium line, (2) the operating line and the equilibrium lines neither touch nor intersect below yNþ1.
 The geometric condition of the minimum liquid flow rate Lmin ; L0min for a given gas flow rate would correspond to the operating line touching the equilibrium line. Thus, the minimum theoretical liquid flow rate (Lmin) is found from the material balance
equatione Lmin ¼ G.(yNþ1  y1)/(xN  xo) and leads to an operating line passing through (xo,y1) and xNþ1 ; yNþ1 , where xNþ1 is the liquid-phase composition in equilibrium with the vapor composition yNþ1. In the construction of Fig. 10.2, this corresponds to xN ¼ 0.0063 and the (Lmin/At) ¼ 147.3 kg mol water/m2. hr, where At is the crosssectional area of the tower. Optimum (economic) operating liquid flow rate (L, L0 ) is usually around 1.5 to 3 times the minimum liquid flow rate. It is important to realize that in the case of the equilibrium line touching or intersecting the operating line at some point P, the gas composition enrichment limit is yP. This point is the “pinch point” for the system, and if yP Rmin ) using Gilliland correlation (Eq. 11. 16)   N  Nmin R  Rmin 0:5124 ¼ 0:7591  0:7532  ; this can also be written as Nþ1 Rþ1   R  Rmin 0:5124 ¼ 0:7591  0:7532  Rþ1

S  Smin S

Say, for R/Rmin ¼ 1.2, R ¼ 1:2  1:947 ¼ 2:336,   S  Sm S  2:34 2:336  1:947 0:5124 ¼ 0:7591  0:7532  ¼ ¼ 0:5087 S 2:336 þ 1 S i.e., S ¼ 2:34=ð1  0:5087Þ ¼ 4:76, the total number of ideal stages in the column. Locating feed tray using Kirkbride correlation (Eq. 11.20) Ratio of number of rectification trays to number of stripping trays is given by "  #0:206 "     #0:206    Sr N r þ 1 xHK;F xLK;B 2 B 0:25 0:3263 2 1367 ¼ ¼  ¼    D 0:35 0:05 521:9 Ss N s þ 1 xLK;F xHK;D ¼ 2:464 Actual number of trays to be provided assuming 70% tray efficiency, Number of rectification trays in column ¼ 4:76  f2:464 =ð1 þ 2:464Þg=0:7 ¼ 4:84, say five trays. Number of stripping trays in column ¼ 4:76  f1  2:464 =ð1 þ 2:464Þg=0:7 ¼ 1:96, say two trays. One tray is additional for the feed entry section. Condenser load, inlet to condenser is vapour at 98.6 C and exit is at 90 C.

11.5 Design illustration e fractionator

319

Based on heat of vaporisation and specific heat of components, estimated Enthalpy of distillate vapour at 98 C ¼ 67,022 kJ/kg mol. Enthalpy of distillate liquid at 90 C ¼ 12,227 kJ/kg mol. Qc ¼ ðR þ 1Þ  D  Enthalpy difference ¼ ð2:336 þ 1Þ  521:9  ð67022  12227Þ ¼ 95401295 kJ=hr ¼ 26500 kW Reboiler load, Qreb ¼ Qc þ Enthalpy leaving with D þ Enthalpy leaving with B  Enthalpy entering with F ¼ 26916 kW Note: Individual enthalpies are calculated based on the problem data. Table 11.8 is worked out for various values of R/Rm and it summarises the effect of R on the number of trays. Corresponding condenser and reboiler loads can be arrived at from the base value at R/Rm ¼ 1.2, i.e., R ¼ 2.3359. The basic relationship is Qc ¼ ðQc Þbase case  ðR þ 1Þ=ðR þ 1Þbase case ;

and

Qreb ¼ Qc þ ðQreb  Qc Þbase

case

Table 11.8 Effect of variation of reflux ratio. Number of actual trays ðh [ 0:7Þ) Total number of trays considering one additional tray for feed section

R/Rm

R

1.05

2.0439

7

3

11

1.10

2.1413

6

3

10

1.15

2.2386

6

3

10

1.20

2.3359

5

2

8

1.25

2.4333

5

2

8

1.30

2.5306

5

2

8

1.35

2.6279

5

2

8

1.40

2.7253

5

2

8

1.45

2.8226

4

2

7

1.50

2.9199

4

2

7

1.55

3.0173

4

2

7

Rectification section

Stripping section

320

Chapter 11 Distillation

The number of trays is close to what was assumed and hence the column pressure drop is reasonably assumed. Also the relative volatility estimation is ok. Column diameter and internals are to be designed based on the content of Chapter 14.

11.6 Flash distillation The process of flash distillation has been introduced right at the beginning of this chapter. This can also be seen as a process of fractionation with a single stage. Since there is only one stage of contact, the purpose of the fractionator column is served by a simple vessel called the flash drum. The feed can be liquid or a mixture of vapour and liquid. Heat may be added, if necessary, to the system in a feed preheat exchanger. In some cases a heating coil or jacket integral with the flash drum is used for adding the heat. Fig. 11.14 shows a typical flash drum with a control valve in the vapour outlet line for throttling the vapour flow to control the drum pressure. Level in the drum is controlled by regulating the liquid outflow rate. We assume that the feed temperature and pressure is sufficient to flash the feed in the drum and generate the vapour and liquid streams of desired composition. Hence, no heat addition arrangements are included.

FIGURE 11.14 Typical Horizontal Flash Drum configuration with associated instrumentation.

Flash distillation problems are usually posed as a separation problem for a feed with known flow rate (F mol/hr), composition (zFi mole fraction of i-th component), temperature (TF) and pressure (PF). Feed has to be flashed to generate the vapour and liquid streams that meet specific composition specification limit either on the vapour or the liquid stream.

11.6 Flash distillation

321

Deliverables from the process design phase forming the basis for mechanical design of a typical flash distillation system are • • • • •

General arrangement of the system, its instrumentation and control e P&ID. Operating temperature (T) and pressure (P) of flash drum. Flow rate and composition of vapour and liquid streams leaving the drum. Feed preheat temperature and preheating load, if preheating is required. Heating coil/jacket details in the drum and the heat load, if required. Drum dimensions, internal fitting like demister pad, vortex breaker and nozzle connections and their locations. Requirement of vacuum, if operation is envisaged under vacuum.

11.6.1 Design equations Design equations related to the process streams Assuming ideal system, the distribution coefficient of component i, psat i ðTÞ P psat is the vapour pressure of the i-th component at temperature T and pressure P. i

P , the liquid-phase activity coefficient gi can be calculated for For nonideal systems: Ki ¼ gi psat i known temperature and composition. These calculations are straightforward and are dealt in details in text books on thermodynamics. In case of low pressure systems, where the pure component vapour pressure is below 2 atm, the vapour pressures for component i can be estimated from Antoine’s equation: Ki ¼ yi =xi ¼


 log10 psat ¼ ai  i

bi ðT þ ci Þ

(11.32)

Antoine constants a, b and c for component i are in corresponding units based on the units of T and psat i . When (V/F) is the fraction of moles of feed leaving the drum as vapour and zFi denotes the mole fraction of the i-th component in feed for an n component system.  n  X zFi  ðKi  1Þ ¼0 (11.33) 1 þ ðV=FÞ  ðKi  1Þ 1 The above equation can be solved for ðV=FÞ using an iterative method. Drum pressure P being between the bubble and dew point pressure Pbub and Pdew is the essential condition for existence of both phases in the drum. Only one positive (V/F) value between 0 and 1 will be obtained as a solution of the aforementioned equation. Once the (V/F) value is found from Eqn. 11.33, the liquid and the vapour compositions can be found from the following equation zFi xi ¼ ; (11.34) 1 þ ðV=FÞð1  Ki Þ where xi and yi ð ¼ xi  Ki Þ are mole fractions of i-th component in the liquid and vapour phase.

322

Chapter 11 Distillation

11.6.2 Design considerations The following considerations are helpful in proceeding with the design • •

•

Lower operating pressures require lower drum temperature and lower heating load. Pressure versus vacuum operation e A slight positive pressure is usually preferred in case of combustible systems in order to avoid/reduce the risk of air ingress into the drum. Air ingress may lead to an explosive mixture and is particularly important while handling hydrocarbons. Vacuum flashing is resorted to when there is a maximum limit of a lighter component concentration in the liquid or when maximum recovery of a lighter component is warranted. Operating under vacuum requires a lower operating temperature and heat load but increases vapour volume which in turn results in larger drum size. Vacuum in most cases is created by a steam ejector. Ejector steam consumption increases sharply with higher vacuum requirement and the steam cost may offset the other advantages of vacuum operation. Unless there is some specific requirement, the operating pressure lower limit is kept w0.1 atm (abs). This level of vacuum is easily achieved by a steam ejector coupled with a barometric condenser. Feed preheat versus heating arrangement integral with the drum - A feed preheat exchanger for liquid feed followed by a properly sized restriction orifice allows efficient exchanger design, avoiding vaporisation in the exchanger. The restriction orifice provides the pressure drop and flashing starts immediately downstream. Therefore, the restriction orifice is located close to the drum as shown in Fig. 11.15.

Restriction Orifice close to drum

Vapour Demister pad

Heating steam Feed

Condensate Liquid

FIGURE 11.15 Typical vertical flash drum configuration.

Heating jacket or coil with condensing steam is usually used when the heat load is low. Since the liquid side heat transfer coefficient in the drum is usually controlling, it is not a very efficient process. •

Yield of desired stream e the split of vapour and liquid depends on the drum temperature and pressure. Maximising the yield of the desired stream is a design objective. The operating temperature and pressure maximising the yield of desired stream is found out by trial. It is easier

11.6 Flash distillation

•

•

•

323

to fix a temperature, compute the corresponding (V/F) splits by varying the drum pressure between the bubble and dew pressure of feed and tabulating the splits that meet the specification requirements. The design operating condition is chosen from the tabulated data, considering the balance between the yield and the feasibility of adopting the corresponding design pressure and temperature. Choice between the horizontal and the vertical drum e Horizontal drums provide a longer travel path and also a larger cross section for vapour travel that reduces vapour velocity. This helps separation of the phases. Vertical drums on the other hand occupy less floor space. Length to diameter ratio of vertical drums is usually between 3 and 5. While sizing a vertical drum, if the length/diameter estimate exceeds 5, a horizontal drum is designed. Demister pads e These are provided when the chance of liquid carryover is to be minimised. Typical demister pads are 80e100 mm thick wire gauge pads and offer below 25 mm water column pressure drop. These are bought out items. In horizontal drums the demister is provided only for a section of the drum in the area below the vapour outlet nozzle. This is schematically shown in Fig. 11.14 and 11.15. One may refer to Chapter 17 for sizing the features of the flash drum.

11.6.3 Design steps The design steps consist of (i) Deciding the design operating conditions e the drum temperature and pressure required to meet the composition specification assuming equilibrium conditions (TV ¼ TL & PV ¼ PL). This step generates the flow rate and composition of the vapour and liquid streams leaving the drum, preheat temperature requirement, heat load through preheat and/or through heating arrangement integral to the drum. (ii) Firming up the general arrangement of the process. This includes arrangement for preheating the feed and/or heating arrangement for the drum, arrangements for controlling the feed rate, drum temperature, pressure and liquid level. This generates the P&ID of the system. (iii) Sizing of the drum is based on applicable design equations and industry practices to ensure safe and trouble-free operation of the system. This includes decision on the drum configuration (orientation) e vertical or horizontal; requirement of mist eliminator/demister pad to minimise liquid entrainment in the vapour stream; sizing of feed inlet and vapour and liquid outlet nozzles; location and sizing of nozzles for drain, vent, safety valve fixing, utility connections; location and sizing of nozzles for instrumentation e temperature, pressure and level measurements. (iv) Detailed mechanical design of the drum and the associated system. Sizing of the flash drum Maximum allowable vapour velocity on the liquid surface is uV;max ¼ SF  fðrL  rV Þ=rL g1=2

(11.35)

where rL and rV are the density of liquid and vapour (kg/m3) and mL ; mV are the corresponding flow rates (kg/sec).

324

Chapter 11 Distillation

The system factor SF for horizontal and vertical vessels (with mist eliminator) is given in Table 11.9 against values of the separation factor s. s ¼ ðmL = mV ÞðrV =rL Þ1=2

(11.36)

Table 11.9 Separation factor at 85% flooding velocity (with mist eliminator). System factor SF(m/sec) s

For vertical drum, SFv

For horizontal drum, SFh

0.006

0.0762

0.0953

0.008

0.0914

0.1142

0.01

0.1006

0.1258

0.02

0.1219

0.1524

0.04

0.1341

0.1676

0.06

0.1341

0.1676

0.08

0.1310

0.1638

0.1

0.1280

0.1600

0.2

0.1128

0.1410

0.4

0.0884

0.1105

0.6

0.0671

0.0839

0.8

0.0549

0.0686

1.0

0.0488

0.0610

2

0.0228

0.0285

4

0.0101

0.0126

6

0.0055

0.0069

A higher value of the system factor SF allows a higher allowable maximum vapour velocity and leads to lower cross section. Since higher disengagement height (hdisengaging) would reduce entrainment and allow higher SF to be used, a second approach to sizing is based on Table 11.10 that tabulates (maximum) allowable SF against values of hdisengaging.

Table 11.10 SF values based on disengaging height (hdisengaging) with mist eliminator. hdisengaging (mm)

75

100

125

150

175

200

225

250

275

300

325

350

Allowable SF value

0.12

0.15

0.19

0.22

0.25

0.29

0.32

0.35

0.38

0.40

0.42

0.43

11.6 Flash distillation

325

Vertical drum The critical dimensions of a vertical drum are shown in Fig. 11.16. Vapour outlet

h1

Feed inlet

DØ

High level of liquid

h2

Normal level Low level

hL

Liquid outlet

FIGURE 11.16 Vertical flash drum with its critical dimensions.

The minimum diameter (dmin ) in m is 1=2
 2 dmin ¼ pffiffiffi  mv = rv: uv;max p

(11.37)

dmin is normally rounded off to the next higher multiple of 10 mm and incremented in steps of 150 mm, if required to be revised. All variables in Eqn. 11.37 have to be in consistent units. Feed inlet nozzle Some amount of feed vaporises in the feed nozzle and hence the same cannot be sized based only on volumetric flow of entering liquid. Inlet nozzle size is based on average (rmix ) density of the vapour and liquid leaving. rmix ¼ ðmL þ mv Þ=ðmL = rL þ mv = rv Þ; kg=m3

(11.38)

Maximum and minimum allowable inlet nozzle velocity are estimated from the following empirical expressions pffiffiffiffiffiffiffiffi umax;nozzle ¼ 122= rmix ; m=sec (11.39) umin;nozzle ¼ 0:6  umax;nozzle ; m=sec

(11.40)

326

Chapter 11 Distillation

Minimum inner diameter of inlet nozzle, di;nozzle is 1=2
pffiffiffi  
di;nozzle ¼ 2 = p  ðmL þ mV Þ= rmix  umax;nozzle ; m

(11.41)

From the standard pipe size chart the nearest higher id pipe is chosen for use as nozzle and its corresponding outer diameter value do;nozzle ðmmÞ is noted. Nozzle locations, for a vertical drum h1 ¼ 900 þ do;nozzle ðmmÞ=2; mm or 1200 mm; whichever is higher h2 ¼ 300 þ do;nozzle ðmmÞ=2; mm or 450 mm; whichever is higher.

(11.42)

Time to fill the drum from minimum to the maximum liquid level is decided based on the emergency response of the operator and this is usually between 3 and 6 min. This is the surge time (ts , sec). Surge volume Vs ¼ ts  ðmL =rL Þ; m3 hL ¼ ð4=pÞ  Vs =d2 ; m hTotal ¼ h1 þ h2 þ hL

(11.43)

hTotal =d ¼ ðh1 þ h2 þ hL Þ=d Desirable hTotal /d is between 3 and 5. In case hTotal /d exceeds 5, a horizontal drum is chosen. Outlet nozzles for vapour and liquid Sizing of liquid and vapour outlet nozzles are based on typical vapour line velocity around 15 m/s and the liquid line velocity between 2 and 3 m/s. Orientation and location of the remaining nozzles are done based on operability consideration, e.g., the drain and the vent are located at the highest and lowest points in the drum, utility connections (plant air and water) are located at accessible positions, etc. Finally, a general arrangement drawing of the drum is prepared that shows all nozzles, fixtures and vessel support. This drawing becomes the basis for the fabrication drawing of the vessel.

11.7 Design illustration e flash distillation Design problem: A 12 m3/hr stream of ortho-xylene is contaminated with 1%w/w benzene and 1.45% w/w toluene. Design a flash distillation process to bring down benzene content in the stream to below 0.5%w/w. Solution The problem states the composition limit on the liquid stream from the drum. This needs to be met while ensuring that the product yield should be as high as possible. This sets the benzene content in the liquid stream to 0.5% w/w. Component properties are tabulated in Table 11.11.

11.7 Design illustration e flash distillation

327

Table 11.11 Properties of the B-T-(O-x) system. Benzene (B)

Toluene (T)

Ortho-xylene (O-x)

Molecular Weight

78

92

106

N.B.P. (K)

353.05

383.6

417.4

Density, r (gm/cc)

0.8787

0.8636

Constants in Antoine’s equation log10(p ) ¼ ab/(T þ c); p sat

sat

in kPa, T in

0.8800 C

a

6.01905

6.08436

6.12699

b

1204.637

1347.620

1476.753

c

220.089

219.787

213.911

Cpl (kJ/kmol K), average

134.8

155.96

187

Cpv (kJ/kmol K), average

82.44

103.7

132.5

Heat of vaporisation at NBP, l (kJ/kmol)

30.77

38.06

36.24

% w/w in feed

1.1

1.4

97.5

Mole% in feed (zF)

0.0149

0.0160

0.9691

l at temperatures T1 and T2 are related by the equation: (l 1/l2) ¼ {(TcT1)/(TcT2)}0.38, where Tc is the critical temperature of the component.

Flash drum operating temperature and pressure Since the temperature at which the feed may be available is not mentioned, it is assumed that the feed is from a tank at ambient temperature, say w32 C. The N.B.P. of the limiting component benzene is 78 C. This is not too high and easily attainable with heating using low pressure steam. Therefore, we adopt drum temperature in the range 70e110 C so that it is achievable after preheating with steam or any other hot stream. For a particular drum temperature (T), the bubble (PBub:F ) and the dew point (PDew:F ) pressure of the feed are calculated from pure component vapour pressures and feed mole fractions as X X
 PBub:F ¼ zFi psat zFi = psat i .& PDew:F ¼ 1= i i¼1;3

i¼1;3

Flashing of feed at a pressure P between the bubble point pressure PBub:F and dew point pressure PDew:F would give rise to vapour and liquid phases coexisting in the drum at temperature T. The V/F value and the corresponding liquid mole fraction (xi ) are to be calculated from Eqs. 11.33 and 11.34. Solution of Eq. 11.33 for V/F requires trial or use of a convergence routine, e.g., NewtoneRaphson. The acceptable P with which the drum can be operated shall be the one with benzene concentration 0.5% w/w or lower in the liquid. It may be noted that selecting an even lower pressure would result in benzene concentration below 0.5%w/w, which is acceptable but this will be at the cost of product

328

Chapter 11 Distillation

stream yield. Therefore, the optimum P corresponding to a T is the one with 0.5% w/w benzene in the liquid phase. It is simpler to compute xi values for different values of P and select the one that has 0.5% w/w benzene in the liquid phase. Therefore, values of T and its corresponding optimum P along with the V/F and phase compositions are estimated and tabulated in Table 11.12.

Table 11.12 Results of flashing at different temperature and corresponding optimum pressure. Liquid-phase mole fractionsa

Vapour-phase mole fractionsa

T (o C)

Flash drum Pressure (kPa)

V/F Mole ratio

Benzene

Toluene

Benzene

Toluene

%w/w Benzene in bottom product

70

9.08

0.1728

0.0067

0.0119

0.0540

0.0357

0.49

82.94

75

11.14

0.1786

0.0067

0.0119

0.0522

0.0349

0.5

82.37

80

13.56

0.1906

0.0067

0.0118

0.0497

0.0339

0.49

81.16

85

16.42

0.1964

0.0067

0.0118

0.0482

0.0332

0.5

80.58

90

19.75

0.202

0.0068

0.0119

0.0468

0.0325

0.5

80.01

95

23.6

0.2146

0.0067

0.0118

0.0447

0.0317

0.5

78.75

100

28.06

0.2203

0.0068

0.0118

0.0435

0.0311

0.5

78.18

105

33.17

0.2334

0.0067

0.0117

0.0416

0.0303

0.5

76.88

110

39.04

0.239

0.0068

0.0117

0.0406

0.0298

0.5

76.31

%w/w Yield of bottom product

a

Ortho-xylene mole fraction to be found by balance.

Finalising design T and P High % yield (82.37e82.94) is observed when the flashing temperature is 70e75 C and flashing pressure is 9e11.5 kPa. We choose design condition of 75 C and 10 kPa as the operating pressure (vacuum) can be obtained by using steam ejector, followed by condenser and a separator for hydrocarbon and water/condensate. Steam consumption increases substantially at higher vacuum (pressure below w10 kPa), thus increasing the ejector size/capacity, condenser size (area) and generates more condensate contaminated with hydrocarbon. One may also consider a condenser cooled by cooling water for the hydrocarbon vapour exiting the drum. Part of the vapour not condensing therein may be sucked by the ejector condenser. This would reduce steam consumption at the cost of additional investment for the condenser and cost for its cooling water consumption. Although a true cost optimisation can be done by considering the heat input requirement through a feed preheater along with other considerations, we adopt T ¼ 75 C and P ¼ 10 kPa as design operating conditions. The process conditions are marked on the schematic diagram in Fig. 11.17.

11.7 Design illustration e flash distillation

329

V 75°C, 10 kPa Steam

F 12 m3/hr

10 kPa 75°C

Condensate L 75°C, 10 kPa

FIGURE 11.17 Schematic showing the design conditions of the flash drum.

Drum design

Feed (F, liquid) Flow rate Component

% w/w

MW

r (mt/m )

m /hr

mt/hr

Flow rate mt mol/hr

B

1.1

78

0.8787

0.132

0.116

0.00149

3

3

T

1.4

92

0.8636

0.173

0.149

0.00162

O-x

97.5

106

0.88

11.695

10.293

0.09710

0.8798

12

10.558

0.10021

100

V/F ¼ 0.1786 mol/mol; V ¼ 0.10021  0.1786 ¼ 0.0179 mt mol/hr.

Product (L, liquid) Component

% mol/mol

B

0.67

mt mol/hr

mt/hr

m3/hr

5.548  104

0.0433

0.0492

4

T

1.19

9.854  10

0.0907

0.1050

O-x

98.14

0.0812698

8.6145

9.7892

100

0.08281

8.7485

9.9434

rL ¼ 8:7485=9:9434 ¼ 0:8798 mt=m3 ¼ 879:8 kg=m3 mL ¼ 8.7485 mt=hr ¼ 8748.5 kg=hr

330

Chapter 11 Distillation

Vapour T ¼ 75 C; P ¼ 10 KPa rVBenzene ¼

78  273  10 ¼ 2:697  104 mt=m3 ¼ 0:2697 kg=m3 22:4  103  ð273 þ 75Þ  101:3

rVToluene ¼

92  273  10 ¼ 3:18  104 mt=m3 ¼ 0:3181 kg=m3 22:4  103  ð273 þ 75Þ  101:3

rVoxylene ¼

106  273  10 ¼ 3:665  104 mt=m3 ¼ 0:3665 kg=m3 22:4  103  ð273 þ 75Þ  101:3



Vapour stream (V) at T [ 75 C; P [ 10 kPa mt mol/hr

mt/hr

m3/hr

rV mt/m3

Benzene

9.352  104

0.0729

270.30

2.697  104

Toluene

6.346  104

0.0584

183.59

3.181  104

0-xylene

0.0158302

1.6782

4578.99

3.665  104

0.0174

1.8095

5032.88

3.595  104

Component

rV ¼ 3.595 x 104 mt=m3 ¼ 0.3595 kg=m3 mV ¼ 1.8095 mt/hr ¼ 1809.5 kg/hr Separator factor (s) s ¼ ðmL =mV Þ  ðrV =rL Þ1=2 ¼ ð8748:5=1809:5Þ  ð0:3595=879:8Þ1=2 ¼ 0:098 From Table 11.9 (system constant, SF, for vertical separator with demister pad @ 85% flooding), SFV ¼ 0:1310 þ

uv;max ¼ SFv 

ð0:1280  0:1310Þ  ð0:098  0:08Þ ¼ 0:1283 0:02

    rl  rv 1=2 879:8  0:3595 1=2 ¼ 0:1283  m=sec ¼ 6:346 m=sec 0:3595 rv

Inlet nozzle rmix ¼

mL þ mV ð8748:5 þ 1809:5Þ  ¼ 2:0935 kg=m3 ¼ 8748:5 1809:5 ðmL =rL þ mV =rV Þ þ 879:8 0:3595

11.7 Design illustration e flash distillation

For inlet nozzle (from Eqs. 11.39 and 11.40) ( 1=2 umax ¼ 122=rmix ¼ ¼

umin

0:6  umax

122=ð2:0935Þ1=2

¼

331

¼ 84:32 m=sec

50:5 m=sec

From Eq. (11.41) ID of nozzle di;nozzle (m) is given by 2 pdi;nozzle

4

 umax ¼

10:558 1  2:0935  103 3600

or 

di;nozzle

10:558  4 ¼ 2:0935  103  3600  p  84:32

1=2 ¼ 0:145 m 00

We choose 600 NB, schedule 40 steel pipe with inner diameter 6.065 ¼ 154 mm and outer diameter 6.62500 ¼ 168 mm as feed nozzle. Drum internal diameter From Eqn. 11.37, dmin , the minimum diameter (m) of the vessel is given by p 2 1:8095 1 dmin  uvap:max ¼  m3 =sec 4 3600 3:595  104 or  dmin ¼

1:8095  4 3600  3:595  104  p  6:346

1=2 ¼ 0:53 m ¼ 530 mm

We choose an additional cushion of w150 mm for diameter and set d ¼ 700 mm ¼ 0.7 m. Deciding vessel height and checking for hT =d limits (Refer Fig. 11.16) hv ¼ h1 þ h2 Feed nozzle od ¼ 168 mm h1 ¼ 900 mm þ od=2 ¼ 900 þ 168=2 ¼ 984 mm Minimum acceptable h1 ¼ 1200 mm Hence, h1 ¼ 1200 mm h2 ¼ 300 mm þ od=2 ¼ 300 þ 168=2 ¼ 384 mm; say 390 mm Minimum acceptable h2 is 450 mm. Hence, h2 ¼ 450 mm Height of vapour space (excluding vessel head) is hv ¼ h1 þ h2 ¼ 1200 þ 450 ¼ 1650 mm. Residence time of liquid is 3e5 min for vapoureliquid separator.

332

Chapter 11 Distillation

We consider 4 min (240 s) 9:9434 3 m ¼ 0:663 m3 3600 This shall correspond to the normal level at the average of high and low level of liquid. We consider a level span of 300 mm, and further considering 300 mm above the high level, we locate the inlet nozzle centre line (see Chapter 17 for additional details). Depth of liquid filled space (excluding vessel bottom dished end) is r Liquid holdup ¼ 240 

hL ¼ p

0:663  ð0:7Þ2

¼ 1:72 m

4 rhT = d ¼ ðhL þ hV Þ=d ¼ ð1:72 þ 1:68Þ=0:7 ¼ 4:86 < 5 Hence, a vertical vessel is acceptable. Vapour and liquid outlet nozzles Typical nozzle velocities for vapour and liquid services are w15 m=sec and 2e3 m/s, respectively. For vapour exit nozzle based on 15 m/s vapour velocity,  p
2 5032:88 dV;nozzle  15 ¼ 4 3600 dV;nozzle ¼ 0:344 m ¼ 344 mm Closest internal diameter standard nozzle (pipe) is 1400 NB, Schedule 40, with inner diameter ¼ 13:37600 ¼ 340 mm f is used as vapour nozzle. For liquid nozzle based on 2.5 m/s velocity, p 2 8:7485 d  2:5 ¼ 4 L;nozzle 3600  1=2 9:9434  4 dL;nozzle ¼ ¼ 0:0375 m ¼ 37:5 mm p  2:5  3600 We provide a minimum nozzle size of 200 NB, Schedule 40(di ¼ 2:06700 ¼ 52:5 mm f) as liquid nozzle. Vessel support and other nozzles for vent, drain, level measurement, temperature sensing and safety valve mounting are also to be provided as per guidelines provided in Chapter 17. One may use the information in Chapter 17 for details on instrumentation and other nozzles.

11.8 Batch distillation Batch distillation, as the name suggests, is a process where the distillation is carried out on batches of feed. In its simplest form, there is only one stage of vapoureliquid contacting and that involves a ‘charge still’ that may have heating coils or jacket integral to the vessel. The vessel is charged with a fixed amount of liquid mixture that is to be separated by boiling and subsequent condensation. In some cases, there can be an external heater through which the still charge is circulated and the heated charge flashes on return to the still. The vapour leaving the still goes to a condenser from where the condensed

11.8 Batch distillation

333

liquid is collected in the accumulator vessel. Since the distillate gradually gets heavier, it can be collected at different time intervals as batches of product with the desired composition. For example, the equilibrium vapour composition corresponding to 50 mol % benzene in liquid (at w1 atm pressure) is about 70 mol %. Therefore, batch distilling a feed mixture of 50:50 mol ratio benzene and toluene can produce distillate batches with 60% benzene, 55% benzene, etc., collected at different time intervals for appropriate durations. Different quality of distillates may be collected in different accumulator vessels. The quality of the distillate is judged from the samples of distillate collected and also from the temperature of the vapour entering the condenser. This is usually a simple system with a heated still and a condenser operating around atmospheric pressure. Such a configuration produces poor separation with either high or low concentrations of light component and involves fairly large energy expenditure. The performance can be improved by operating the still at a lower pressure but an operating pressure (Pop) lower than 0.1 kg/cm2 (abs) significantly increases the ejector steam consumption for creating vacuum and the option becomes uneconomical. Redistillation of distillate in subsequent batches can also be performed to obtain higher distillate purity. On a different note, in small-scale batch distillation set ups with only the still and condenser, equivalent number of vapoureliquid contacting stages is slightly more than one. This happens due to the heat loss from the rising vapour through the equipment wall and its partial condensation providing reflux that aids the separation. The ASTM D86 distillation test apparatus for light petroleum cuts thus provides separation closely equivalent to 1.1 Batch distillation with multiple contacting stages theoretical stages. The vapour condensing in the neck of the standard ASTM D86 flask creates the internal liquid reflux. When the required purity of the product is higher than that attainable with a single stage of contacting or for systems with low relative volatility ða < 3Þ, batch distillation is carried out with multiple vapoureliquid contacting stages. In this case (Fig. 11.18) the vapour rises through a packed or tray tower where it meets a counter flowing liquid reflux. Batch distillation columns most commonly employ sieve trays without downcomer or packed beds. As only a few trays are usually required, low tray efficiency is not of much concern and sufficient extra trays are provided. Employing higher efficiency trays, e.g., cross flow sieve trays, bubble caps, etc., cost much more for such small scale applications. The tower is usually mounted directly on the charge still (Fig. 11.18A) in case of small units. In other cases the tower is placed next to the still at an appropriate elevation to allow gravity flow of liquid from the tower bottom back to the still (Fig. 11.18B). The vapour leaving the tower top enters the condenser through the overhead vapour line. Part of the condensed overhead stream is diverted back to the column as reflux and the rest is collected in the accumulator vessel. The reflux usually is a metered flow, typically through a rotameter. In larger plants, a pump may be used for sending the reflux from the accumulator to the tower top; else it is a gravity flow, requiring the condenser and the accumulator vessel to be located at an appropriate elevation. Pressure control in batch distillation is similar to a rectification column and is achieved by throttling the noncondensable vapour venting from the accumulator top. In case the tower is to be operated under

334

Chapter 11 Distillation

(A)

Cooling water in

Cooling water out Condenser

EQUALISER line to condenser shell

Reflex line

(B)

Condenser Cooling Water 3-way valve to control reflux Distillate

Tray column Distillate

Vapour

Packed column

Steam Change still

Condensate Liquid

Reboiler Bottoms

Steam

Condensate

Bottoms

FIGURE 11.18 Multistage Batch Distillation (A) Tray column above still, (B) Packed column beside still.

vacuum, the vent is connected often through a surface condenser to the suction of an ‘ejectore barometric condenser’ arrangement. The surface condenser upstream of the ejector condenses much of the overhead vapour and reduces the vapour load to the ejector. The residual streams (bottoms) is usually taken out of the still after the dilatation ends once the still is cooler and safer to drain. Instrumentation scheme Instrumentation of the process consists of temperature measurements of the liquid in the still, vapour entering the condenser and the liquid collected in the accumulator vessel. Pressure is measured at the top of the still. This needs to be a compound gauge measuring both vacuum and pressure as condensation of vapour after shutdown may lead to vacuum in the system. The accumulator and the charge still are provided with level gauge glasses. In small plants these instruments locally indicate the readings, whereas in large plants there may be telemetry arrangement along with alarms associated with the process parameters. The charge still heat input should be regulated as this decides the start-up period and also production rate of distillate. In case of electrical heating the power to the heating element is regulated. Opening of the control valve in the heating steam supply line is manipulated to control heat input in case of steam heated still. Sudden increase of input heat may lead to large amount of vaporisation in the still and carryover of liquid along with distillate vapour. This leads to poorer

11.8 Batch distillation

335

fractionation and needs to be avoided and a steady regulated heating rate in the still is achieved by the flow control of the circulating heating oil or by pressure control of the heating steam supply. Batch operation Batch distillation is an unsteady state operation, i.e., the distillate and still compositions continuously vary with time. The process operates through time cycles of i. loading of the feed charge in the batch still ii. heating up/start-up phase e when the process absorbs energy without any distillate being withdrawn. (Duration: t1) iii. production run e this is the period during which the distillate is withdrawn. (Duration: t2) iv. cooling down the still and preparing the equipment for the next batch. (Duration: t3) Operation without fractionating stages and reflux e single-stage batch distillation Following the initial inspection of the readiness of equipment the charge is loaded into the still. The coolant flow in the condenser, which is cooling water in most cases, is started. If the process needs to be connected to a vacuum system, the same is done and vacuum pulling is done right at the cold state. If vacuum is pulled later when the still is already heated up, sudden vaporisation and foaming lead to carryover of liquid with the vapour. Production run starts as the first drop of liquid falls into the accumulator vessel. During production run the distillate cut is collected. In case of more than one cut, separate accumulators are used. The still pressure, if required to be kept at slightly above atmospheric pressure, is regulated by controlled bleeding of noncondensable gases from the accumulator vessel top. In small set ups there may be a water seal through which this gas may be bled, limiting the accumulator gauge pressure to the seal liquid dip head. Such a seal system should also include arrangement to prevent seal-liquid suck-back in case of vacuum creation in the equipment during cooling down or even otherwise. This is usually ensured by a liquid trap installed in the line upstream of the seal. The cut specification control is done by (i) periodic sampling of the liquid from the accumulator and its analysis, (ii) ascertaining the yield of the cut against previous instructions based on assessment of feed composition. The vapour temperature leaving the still is the dew point corresponding to the instantaneous composition of the vapour. This temperature is monitored to control the components going to a particular cut of distillate. Operation of multistage column The operational procedures of start-up, production and shutdown phases may differ slightly for different column configurations but in general, the operation follows the following sequence. Start-up phase In practice, an empty conventional multistage batch column is started up in the following sequential steps: 1. The still is charged with the feed to be processed and heat is applied to bring the material to its boiling (bubble) point temperature. 2. Depending on the heat input, a part of the charge vaporises and the vapour travels upward through the column internals. Part of the vapour condenses in contact with the cold column and its internals that get heated in turn. The condensate travels down the column to the still as reflux. As the tower and internals get hotter, the condensing amount reduces and after a while the vapour reaches the condenser.

336

Chapter 11 Distillation

3. At this time, the coolant flow to the condenser is started and the liquid condensed starts accumulating in the reflux drum. Some product may also be collected during this period. The reflux valve is opened once the liquid level in the reflux drum reaches the normal operating level. 4. If no product is withdrawn in step 3, the column is run under total reflux operation, i.e., the entire vapour reaching the condenser is condensed and refluxed back. This is continued till a steady state, evidenced by steady temperatures in the column top and the still, is reached. Alternatively the operation can be continued till the distillate composition reaches the desired product purity. Duration of the first step is usually small compared to the overall batch time, whereas the duration of steps 2e4 is important and in some cases it may take a long time to reach a steady state or the desired initial distillate composition. Production phase Production period starts with withdrawal of distillate product from the reflux drum. Operation in this period and its duration depends on the required specifications of the product or on the economics of the process. Operation in the production phase can be under the following conditions: (i) Constant distillate composition operation (variable reflux operation): On reaching the required distillate purity, the start-up period ends. Product take off is started and a constant composition product is collected by steadily increasing the reflux ratio until a specified amount of distillate has been collected. During this period the column top temperature is kept constant by continuously increasing the reflux flow. The operation ends when the reflux ratio has attained some high value considered to be uneconomic. (ii) Constant reflux operation: The total reflux start-up period ends with the column reaching its steady state, i.e., the still and the column top temperatures become steady. During this production phase, the column operates with a suitably preset fixed reflux ratio so that the distillate purity from the initial to the end reduces from a high value to a lower value and the collected distillate with average purity matches the target. In general, the constant reflux operation is preferred. The pros and cons of each operation is discussed in Section 11.8.1. (iii) Optimal operation: A third mode of operation is a trade-off between the above two modes. An optimal reflux policy is chosen so that some objective function is optimised (minimum batch time, maximum product yield, maximum profit, etc.), subject to constraints (e.g., on product amount and purity) at the end of the process. Shutdown At the end of the production phase, a batch distillation column is shut down in the following sequence: 1. Heat supply to the column is cut off. 2. Holdup in the column is collected in the still. 3. Condenser holdup may be mixed with the top product or with the still material.

11.8 Batch distillation

337

11.8.1 Design Choice of the mode of operation is the first step in designing a batch distillation process. Single-stage distillation leads to maximum recovery of the light component but the concentration of the same in the distillate is low. Due to the interaction of the rising vapour with the liquid reflux falling down the column, the rate of change of distillate composition is much slower when the operation is with multiple stages and reflux. Operation with refluxing requires more energy compared to ‘no reflux’, i.e., singlestage operation. Operation with varying reflux ratio needs highest energy requirement though the latter permits highest purity of the overhead product. The reflux flow rate (and in turn the reflux ratio) is manipulated to control the temperature of the vapour to the condenser (and in turn the distillate composition). The highest fraction of more volatile is also left in the still for varying reflux ratio. Constant reflux ratio operation is preferred due to its operational simplicity and the operation with varying reflux ratio to obtain constant distillate composition is selected when the quality requirement of distillate is very strict. Batch distillation design problems are usually posed as the feed and product specifications being specified, along with feed processing rate in terms of quantity per day. Operational constraints like single/two/three shift operation of the plant may also be specified based on which the size and the number of batches per day may be decided. The feed charge volume per batch decides the still size (volume) and is considered while deciding in favour of batch distillation. The batch final volume left in the still should not be too small so that its practical design considering submergence of the heating surface is difficult. Also it should not be too large. The designer also decides whether a column with contacting stages is needed in addition to the vapoureliquid contacting in the still. As already mentioned, a column is required when the difference in purity of the distillate and the feed is large and attainment of the desired product purity requires two or more stages. A McCabeeThiele diagram drawn for the rectification section provides information to design the column. The minimum number of stages is estimated under total reflux conditions with bottom product composition xB ¼ xF and top product composition xD . In order to draw a finite quantity of top product, the reflux ratio needs to be decreased. Alternatively the x-y plot can be altered by selecting a lower operating pressure. However, this is rarely adopted from cost and safety considerations as discussed in Section 11.2. The details of the design procedure for a simple batch with no reflux and a batch with column for both constant reflux and constant top product composition are provided in the following section. Design equations Binary system with no reflux: The design for a binary system with no reflux is based on Rayleigh equation relating the liquid amount (moles) and its composition (mole fraction) in the still at any instant of time with the instantaneous composition of the vapour generated during differential distillation. B ln ¼ F

ZxB

dx y  x

(11.44)

xF

where B and xB are the amount (in moles) and mole fraction of the more volatile component in the still at the end of distillation and F and xF are the moles and corresponding composition of the feed (initial charge). y is the vapour composition in equilibrium with x at any instant of time. For ideal

338

Chapter 11 Distillation

mixtures, y is related to x using relative volatility a, which is often temperature dependent. The relationship already presented earlier in Eq. 11.11 is ax y ¼ 1 þ ða  1Þx Substituting the above expression in Eq. 11.44 gives  1=ða1Þ   B xB 1  xF a=ða1Þ ¼  F xF 1  xB

(11.45)

Since the distillate becomes progressively less rich in the lighter component, one would be interested in the mean distillate composition (xDavg ) given by xF  xB (11.46) xDavg ¼ xB þ 1  ðB=FÞ Usually the design inputs are initial charge/Feed (F) and its composition (xF ) and either the distillate (xDavg ) or the still composition (xB ) at the end. The corresponding outputs are either (xB ) and B (bottom product composition and amount) or xDavg and D. Multistage batch distillation of binary system with reflux: The calculation of multistage batch distillation of binary mixtures is also based on Rayleigh equation with y replaced by the distillate mole fraction xD B ln ¼ F

ZxB xF

dxB xD  xB

(11.47)

The concentration (xD ) of the distillate depends on vapoureliquid equilibrium or a, separation efficiency of the column (the number of equilibrium stages N or the number of transfer units NTOG) and the reflux ratio R. Mathematically, xD ¼ f ðxB ; a; N; RÞ

(11.48)

The functional form of Eq. 11.48 depends on the operation, i.e., constant reflux ratio, constant distillate composition or mixed mode. The reboil duty (Q) and the corresponding design rating would also be different for each mode of operation. The option selected would dictate the number of theoretical stages (N or NTOG) required for separation. The corresponding vapour and liquid traffic also get fixed. Calculation of height and diameter of the column additionally require details of the internals (packing, trays, etc.). (i) Constant reflux ratio operation: The top product purity in this mode decreases as distillation proceeds (Fig. 11.19A) and xDavg is the targeted distillate product composition. Therefore, initial xD has to be higher than the targeted xDavg and distillation would continue even after the instantaneous xD drops below xDavg. Usually the inputs are N, R and xF and the outputs are D, xDavg and energy requirement for distillation (Q). In order to evaluate the aforementioned, a distillate composition (xD1 ) is set and an operating line is drawn followed by the usual determination of the concentration (xB1 ) in the still by stepping off the

11.8 Batch distillation

339

number of stages (N). Subsequently, another distillate composition (xD2 ) and the corresponding bottom concentration (xB2 ) is determined similarly.
R The operating lines are parallel (Fig. 11.19A) . In this way, the relation between xD and xB is as the reflux ratio R is constant and the slope is Rþ1 determined over the operating range and the integral in Eq. 11.47 is solved by evaluating the area
1  under the curve plotted with xD x on the ordinate and the actual bottoms composition (xB ) on the B abscissa. The area (A) between xF to xB represents the solution to the integral and gives the value of B. The moles of distillate obtained is D ¼ F  ð1  B = FÞ ¼ F  f1  expðAÞg

(11.49)

and the mean distillate concentration (xDavg ) is obtained from the material balance as xDavg ¼

xF  xB  ðB=FÞ 1  ðB=FÞ

(11.50)

The amount of vapour (V) generated in the still is V ¼ D  ðR þ 1Þ

(11.51)

Usually the boil-up rate associated with a specific (distillation) system is known from past experience. In case of a new system, the boils up rate is found by dividing the total vapour load by the time of distillation for only the production run (step iii in batch operation) and the charging time, heat up time, cooling time and clean up time are not included. The heat (Q) required for the separation is related to the other variables as Q ¼ ð1  B = FÞ  ðR þ 1Þ F  lF

(11.52)

where lF is the latent heat of vaporisation of the mixture. (ii) Constant distillate composition operation: In this case the decrease in distillate concentration (xD ) is avoided by continuously increasing the reflux ratio (R) with time. The relationship between still concentration (xB ) and the corresponding reflux ratio (R) is determined graphically on the x-y plot as shown in Fig. 11.19B. xD is located on the 45 degrees line and an initial reflux ratio (R) and the position of the operating line is specified to satisfy N stages and the corresponding still concentration (xB ) is noted. In this case, direct integration of Rayleigh’s equation can be performed to give xD  xF B ¼ F (11.53) xD  xB xF  xB D ¼ F (11.54) xD  xB Correspondingly, the energy requirement (Q) which depends on the variable R is ) ZxB ( Q ðR þ 1Þ ¼ ðxD  xF Þ  dxB F  lF ðxD  xB Þ2 xF

(11.55)

340

Chapter 11 Distillation

The term on the right-hand side of Eq. 11.55 is integrated graphically by plotting the still concentration xB on the abscissa and the term R þ 1 2 on the ordinate. The area (A) under the curve ðxD xB Þ

between xF and xB gives the value of the integral. Thus, the heat requirement Q is found from Q ¼ ðxD  xF Þ  A (11.56) F  lF It is evident from Eqs. 11.52 and 11.55 that the two operating modes require different quantities of heat Q to obtain the same D and xD . The optimum energy requirement is expected to be a combination of the two modes of operation, i.e., combination of distillate concentration decrease and increase in the reflux ratio.

(B) 0.8

(A)

0.7

1.0

0.6

0.8

1

(0.8, 0.5)

0.5 3

y→

y→

3 4

0.4

1 3 4

0.0

(0.4,0.4)

0.6

x→

2'

3'

R = 2.8

0.2

xB = 0.067 xF = 0.25 0.4

R = 0.70

0.4 0.3

2

0.2

(0.512, 0.572)

2

2

0.6

0.2

1,1'

0.16 0.1 0.8

1.0

0.25 0 0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

x→

FIGURE 11.19 McCabeeThiele construction for (A) constant reflux ratio; (B) constant distillate composition.

Multistage batch distillation resembles a fractionator with only the rectification section and is usually designed assuming negligible holdup in the column and condenser as compared to the receiver and kettle. The liquid holdup always present in the column can greatly influence both the yield and purity of the products. Column hold up affects the yield by hindering the sharp decrease of distillate composition and the off specification intermediate fractions are large. The foregoing analysis ignores the holdup in the column and the condenser compared to the holdup in the still. Case of ternary mixtures Rayleigh equation can also be applied to multicomponent batch distillation. For component i, yi ¼ Ki xi which gives dLðKi xi Þ ¼ dðLxi Þ

(11.57)

Since the amount of liquid L in the still and its concentration (xi ) change during operation, it is difficult to evaluate the term dðLxi Þ. In this case, it is advantageous to use component amount (li )

11.8 Batch distillation

341

instead of total liquid amount (L) where the component amount (li ) is defined as li ¼ Lxi , and the liquid concentrations replaced by xi ¼ ðli =LÞ This gives

dL 1 dli 1 dlk ¼ ¼ L Ki li Kk lk

(11.58)

For ideal systems with constant relative volatility (aLK ) with reference to the heaviest component k, dli dlk ¼ alk li lk On integration  1=aik li ¼ lfi

lj lfj

!1=ajk

  lk ¼ lfk

(11.59)

where the amount of liquid in the still is L ¼ Sðli Þ, and the concentration of each component in the still is xi ¼ ðli =LÞ. The same approach can be applied to multistage distillation of ternary/multicomponent mixtures but distillate composition (xDi ) replaces equilibrium composition (yi) (similar to multistage distillation of binary system) to give dB dxBi dxBk ¼ ¼ B xDi  xBi xDk  xBk

(11.60)

In general, determination of distillate composition (xDi ) requires rigorous column simulation and the results are reliable only if the liquid holdup in the column is much lower than the liquid holdup in the still. Batch distillation of ternary (or multicomponent) mixtures is preferentially performed with constant reflux ratio although operation with periodically constant distillate concentration is also possible.

11.8.2 Design deliverables a) Configuration e e e e e e e

Single stage or multistage e number of stages Refluxing arrangement and reflux ratio e constant reflux/constant reflux ratio or constant composition of top product Operating pressure Initial and final temperature of column top and still Initial charge and final liquid holdup volume in still Collection of top product in receivers in case of multiple grades of top product Heat load for still heating e peak, average, variation with time

b) Equipment dimensions Charge tank/still size Volume: The still is chosen to be vertical or horizontal based on considerations same as in flash distillation. This is already elaborated in Section 11.6.2.

342

Chapter 11 Distillation

Total volume of the still ¼ charge volume þ vapour space, typically 30% of total volume. Foaming systems are provided with additional volume usually based on experience. Dimensions: In case of stills with internal heating, the length and diameter of the still are selected such that the heating element remains conveniently submerged at end of operation. This requires a prior estimate of the heating surface and the heating element (coil or tubes) assembly size. For systems with external heater, a circulating pump with adequate head needs to be located appropriately so that the NPSH available is sufficient. In case of single-stage separation, the vessel must also accommodate a demister pad typically of SS wire mesh around 100 mm thick to minimise entrainment. The nozzles are sized based on the same considerations mentioned in Section 11.6.3. Fractionation column Internals e packing/trays: Selection is based on the considerations in Chapter 14, Sections 14.2 & 14.4. Dimensions: Column height has to be sufficient to accommodate packed section, packing support grid and packing restrainer grid (top) or the desired number of trays (usually sieve tray without downcomer) as the case may be. Column diameter is decided from flooding considerations which is a function of flow rates and transport properties of vapour and liquid. The standard design procedure and correlations for packed column flooding and the details of tray sizing are mentioned in Chapter 14. The vapour and liquid flow rates corresponding to which sizing is done and checked depend on the duration of distillation operation. The sizing is checked for i) initial phase where vapour density is minimum ii) end phase where the temperature is maximum iii) maximum reflux ratio Condenser and Heater (internal to the still or external) e These are designed as heat exchange equipment. The external heater may be a shell and tube heat exchanger whose details are provided in Chapter 4. Estimation of heat duty for the condenser and the still heater In either scheme of fixed reflux ratio or fixed top product composition, the total heat supply to the still (Qstill) needs to be decided based on the batch time available. The duty should include an initial load to bring the feed charge and the associated hardware to its bubble point (TbubF), (sensible heat) that should also include the estimated heat loss from the still and the column (Q1) and heat required (Q) during production run, i.e., boil up from (TbubF) to bubble point of bottoms product (TbubB). If facility is kept to use a coolant after stopping of the heat input to the still and cool the final holdup (B) of the still from TbubB to TB, the design of the heat transfer element should also consider this load Q2 as well in its sizing. Qstill ¼ Q1 þQ

(11.61)

Q1 ¼ FCpF ðTbubF  TF Þ þ mCp ðTbubF  TF Þ

(11.62)

where In the above equation the second term is associated with hardware heating and heat losses are considered negligible.

11.8 Batch distillation

Zt Q¼



d ðBSxBi li Þdt dt

343

(11.63)

0

A rigorous evaluation of Q should make use of Eq. 11.63 and for quicker estimate Eq. 11.52 or Eq. 11.55 may be used depending upon the mode of operation selected. Q2 ¼ BCpB ðTbubB  TB Þ þ mCp ðTbubB  TB Þ

(11.64)

where TB is the final temperature of the bottoms product. The average cooling water requirement;

mCW ¼

Q2 m3 =hr t3 CpCW DTcooling

(11.65)

In the above equation t3 ; CpCW and DTcooling are the still cooling time duration, specific heat of cooling water and temperature range of cooling water, respectively. Heat transfer coefficient in each stage of operation is different. Exchanger design must be capable of handling each of these functions. One may also like to operate by selecting a fixed Qstill based on the still heater heat load capacity/ rating. This is particularly true when electric reboilers with fixed wattage are being used. In such case the reflux ratio needs to be suitably altered as distillation proceeds. The heating could be by steam pipes or electrical heating elements which need to remain submerged till the end of the batch. In case of electrical heating, extra heating elements are included. Say if a bench scale set up requires a total power of 12 kW that can be served by using 2 kW elements, instead of using six elements, an additional two elements are usually provided. This requires selection of vessel dimensions such that the liquid holdup at the end of the batch is sufficient to keep all the eight elements completely submerged with a minimum of 80 mm liquid level above and below the heater section. The condenser duty (QC) is estimated based on the vapour leaving the top of the column and is given by Zt   dD QC ¼ ðhV  hL Þdt dt

(11.66)

0

where hV and hL are the enthalpy values of D at its dew point and bubble point temperature, respectively.

11.8.3 Design steps The design steps and calculation procedure are outlined as follows: 1) Decide the operating pressure (Pop ) e Distillation pressure is decided with the same considerations as design of rectification columns that are detailed in Section 11.4.1. 2) Generate the T-x-y data and x-y plot or find a from vapour pressure data obtained from Antoine equation or available experimental data. 3) Note the bubble point at xF and dew point at xD . Find yF corresponding to xF .

344

Chapter 11 Distillation

If yF is reasonably more than xD , a single-stage distillation is sufficient. The calculation can follow the Rayleigh’s equation (Eq. 11.44). In case yF < xD , reduce Pop to increase yF and check if yF is reasonably more than xD . If Pop needs to be reduced below 0.1 kg/cm2 (abs), then distillation under vacuum will possibly turn out to be uneconomic in moderate to large scale operation and more than one stage of fractionation is required. The stages need to be provided by a packed bed or a plate contactor as shown in Fig. 11.18. 4) Initial operation: Starting from empty overhead receiver, filling under no reflux till the desired level in receiver is achieved. Time taken, ti ¼ Di =V , where V is the boil up rate decided considering the production capacity of the plant. Compute xBi bottoms composition at end of receiver filling from     FxF Fð1  xF Þ log ¼ a  log (11.67) Bi ð1  xBi Þ Bi xBi and, from component material balance, Di ¼

FðxF  xBi Þ Bi ðxDi  xBi Þ

(11.68)

or Bi ln ¼ F

ZxB xF

1 dxB ¼ a1 xDi  xB

ln

ð1  xF ÞxBi xF ð1  xBi Þ

þ ln

ð1  xF Þ ð1  xBi Þ

(11.69)

Outputs: Bi , Di , xBi , xDi ; axBi xDi is in equilibrium with xBi , i.e., xDi ¼ 1þða1Þx Bi 5) Selection of the mode of operation viz. (a) constant reflux ratio, (b) constant top product composition (varying reflux ratio) a) Operation with constant reflux Inputs: Bi , Di , xBi , xDi (i) Find Nmin from

 log Nmin þ 1 ¼

   xDi 1  xBi  1  xDi xBi log a

(11.70)

and check N > Nmin , for N that has been decided a priori. N can also be computed from Smoker’s equation, not included here.

11.9 Design illustration e batch distillation

345

(ii) To find Rmin , draw operating line from the point (xDi , xDi ) through the point (xBi ; yBi ) on equilibrium curve to locate the point ðyint ; 0Þ on the ordinate. Calculate Rmin ¼ ðxD =yint Þ  1. (iii) Select R in the range 1.5 Rmin < R < 10 Rmin . Larger the R, higher is the vapour boil-up rate and more is the column diameter and an economic balance may be struck. (iv) Estimate xDavg , V and Q from Eqs. 11.50e11.52 by replacing F with Bi . The expressions are xDavg ¼ ZxB

Bi ¼ ln B

xBi

xBi  xB ðB=Bi Þ ; V ¼ DðR þ 1Þ 1  ðB=Bi Þ

dxB Q ; ¼ ð1  Bi = FÞðR þ 1Þ xDi  xBi FlF

(b) Operation with constant distillate composition (by continuously increasing R) (i) Initially R is minimum corresponding to liquid composition xBi in still. This is found by trial e drawing operating lines through (xDi xDi ) intersecting the equilibrium curve and meeting the ordinate at (yint , 0). On each operating line N stages are constructed starting from (xDi xDi ) and ending in still composition of xB . The operating line corresponding to still composition of xBi is chosen and the corresponding yint value is read off the ordinate scale. Calculate R ¼ ðxD =yint Þ  1 (ii) Find B by integrating Eq. 11.47 after replacing F with Bi B ln ¼ Bi

ZxB xF

dxB xD  xBi

(v) Use Eq. 11.53 through 11.55 to find B, D, Q, and V after replacing F with Bi. The expressions are xD  xBi xBi  xB ; D ¼ Bi xD  xB xD  xB Z xB Rþ1 ¼ ðxD  xBi Þ dxB 2 ðx xBi D  xBi Þ

B ¼ Bi Q FlF and

Z V ¼ Bi ðxD  xBi Þ

xB

xBi

dxB  ðxD  xBi Þ2 1 

R Rþ1



11.9 Design illustration e batch distillation Few drums containing ortho-xylene contaminated with toluene (25% mol/mol toluene, 75% mol/mol Oxylene) are available. As much as possible amount of toluene is to be recovered from this material. Decide on process to recover the toluene. Minimum acceptable purity of toluene-rich phase produced is 40% but this fetches a low selling value. The price improves substantially if the purity is above 55 mole%.

346

Chapter 11 Distillation

Solution 1A: Separation using a simple batch still Since the charge is small, a batch distillation process is considered. The process involves inflammable hydrocarbon and so the operating pressure is set to slightly above atmospheric (¼ 1.1  101.325 ¼ 111.5 kPa) to avoid ingress of air. P ¼ 111:5 kPa; xD;avg ¼ 0:4; xF ¼ 0:25 Checking the feasibility of using a simple batch still with no column and reflux Bubble point of the feed calculated at P based on vapour pressure data of the components, gives Tbub_feed ¼ 136.5 C. Pure component vapour pressures for toluene and O-xylene at 136.5 C are 200.4 and 81.77 kPa, respectively. Relative volatility, a ¼ 200:4=81:77 ¼ 2:45 a  xF Initial composition of vapour leaving the still ¼ ¼ 0:449 w ¼ 0:45. 1 þ ða  1Þ  xF Initial composition is higher than 0.4 and hence it is possible to continue the process and obtain distillate with average composition xD;avg ¼0.4. Based on Eqs. 11.45 & 11.46:  1=ða1Þ   B xB 1  xF a=ða1Þ xF  ðB=FÞ  xB ¼ & xD;avg ¼ F xF 1  xB ð1  B=FÞ

Case 1

Case 2

Case 3 (Desired operation)

Case 4

Case 5

Case 6

xB

0.225

0.200

0.175

0.150

0.125

0.100

ðB=FÞ

0.88

0.79

0.67

0.57

0.48

0.39

xD;avg

0.433

0.416

0.400

0.382

0.364

0.346

0.4157

0.3799

0.3420

0.3019

0.2590

0.2140

xD ¼

a  xB 1 þ xB  ða  1Þ

The desired operation is for xD;avg ¼ 0.4. At start of distillation, xD ¼ 0.45, equilibrium composition w.r.t. xB ¼ 0:25. At end of distillation, xD ¼ 0.56, equilibrium composition w.r.t. xB ¼ 0:3420. Checking for a value at end of operation: Bubble point of B calculated at P at end of operation, gives Tbub_end ¼ 139.7 C and corresponding value of a ¼ 216:6=89:29 ¼ 2:43. This is not much different from the value (2.45) used and no revised calculation is required. Distillate collection starts at still temperature of 136.5 C and ends at 139.7 C. Recovery of toluene ¼ 100 

0:25  ðB=FÞ  xB 0:25  0:67  0:175 ¼ 53:1% ¼ 100  0:25 ð1  B=FÞ  xD;avg

Design of the physical system This requires information on the initial charge volume and temperature. Still: A steam coil heated still is envisaged. Its design involves estimation of the charge volume. The estimated diameter and height must ensure minimum 150 mm liquid depth above the heating coils at end of operation. The vessel design may be guided by available information in Chapter 17. Heating coil details are to be based on heat transfer rate with sufficient extra cushion for which Chapter 2 may be referred to.

11.9 Design illustration e batch distillation

347

Condenser: The condenser is a small shell and tube condenser, cooled with water. Its design can be carried out using the information in Chapter 4. Comment: The recovery of 53.1% toluene is rather low, that too with 40% toluene purity. It is therefore desirable to investigate the option of batch distillation with stages and reflux with around 50% toluene purity. This is illustrated in the following section. Solution 1B: Batch distillation with constant reflux ratio The problem is attempted considering batch distillation with stages and constant reflux ratio. Small-scale distillation would usually employ packed tower for vapoureliquid contacting. Three theoretical contacting stages in the tower (packed bed) and one stage for the reboiler still is considered. The x-y diagram for McCabeeThiele construction is drawn for a ¼ 2.45. For xF ¼ 0.25, by trial and error, the operating line is constructed with total four ideal stages and initial distillate xD ¼ 0.8. This line passes through (0.8,0.8) and has y-intercept of 0.211[ ¼ 0.8/(R þ 1)]. Hence, R ¼ 2.8. Since the reflux ratio R remains constant throughout the operation, the operating lines are parallel to each other. xD reduces from 0.8 as distillation proceeds. Additional operating lines parallel to the original are drawn arbitrarily that pass through (0.7,0.7), (0.6,0.6), (0.5,0.5), (0.4,0.4) and (0.35,0.35). In fact it is easy to draw these lines passing through ðxD ; xD Þ and y-intercept values  xD xD  ¼ can be calculated. ¼ Rþ1 2:8 þ 1 For each operating line starting at (xD ; xD ), the xB that would result for four ideal contacting stages are noted. The value noted corresponding to (0.4,0.4) is ¼ 0.08. The operating lines through (0.8,0.8) and (0.4,0.4) are only shown in Fig. 11.20, else the figure would get clumsy. In fact it is desirable to draw each operating line passing through (xD , xD ) on separate x-y diagram and note the corresponding values of xD and xB . The xD and xB data from the graph is xD

0.8

0.7

0.6

0.5

0.4

0.35

xB

0.25

0.184

0.140

0.107

0.080

0.067

1 ðxD  xB Þ

1.8182

1.9380

2.1739

2.5445

3.1250

3.5336

1.0 0.8 1

0.6

(0.8, 0.8)

y→

2 3 4

0.4

1

0.2

3 4

(0.4,0.4)

2

xB = 0.067 xF = 0.25

0.0

0.2

0.4

0.6

0.8

1.0

x→

FIGURE 11.20 Construction of operating lines for constant reflux ratio operation.

348

Chapter 11 Distillation

A curve of 1/(xD  xB ) versus xB is drawn in Fig. 11.21 to tentatively calculate the area under the curve from xB to xF . As a first trial, xB ¼ 0.067 is considered. 3.6 3.4 3.2

1/ (xD – xB)

3 2.8 2.6 2.4 2.2 2 1.8 0.06 0.08 0.1 0.12 0.14 0.16 0.18 0.2 0.22 0.24 0.26 xB

FIGURE 11.21 Plot of 1/(xD  xB ) versus xB

Z

xF

dxB ¼ ðx  xB Þ D xB Hence, (F/B) ¼ exp(0.4493), i.e., B/F ¼ 0.6381 and

Z

0:25

The area under the curve from the graph

xD;avg ¼

0:067

dxB ¼ 0:4493 ðxD  xB Þ

xF  ðB=FÞ  xB 0:25  0:6381  0:067 ¼ 0:572, highly acceptable as a product. ¼ 1  0:6381 1  ðB=FÞ Toluene recovery ¼ 100  ð0:25  0:6381  0:067Þ=0:25 ¼ 82:9%

For the distillate; lavg ¼ 0:572  38:06e3 þ ð1  0:572Þ  36.24e3 ¼ 37:28e3 kJ=kg mol Total heat required for boil up ¼ ðR þ 1Þ  ð1  ðB = FÞÞ  lavg ¼ 3:8  ð1  0:6381Þ  37:28e3 ¼ 51268 kJ=kg mol charge Therefore, it is feasible to have a batch distillation unit with three ideal stages in a column, operating with constant reflux ratio 2.8, and recover about 82.9% toluene with 57.2% mol purity. However, during operation the trays or packed tower will have a hold up that would drain into the still increasing toluene content in the final hold up in the still. This may slightly reduce the recovery percentage. In any case with this basic design scheme, 80% recovery of toluene with 55% purity can safely be committed.

11.9 Design illustration e batch distillation

349

Solution 1C: Batch distillation with constant distillate composition The problem is attempted considering batch distillation with stages and constant distillate composition. Three theoretical contacting stages in the tower (packed bed) and one stage for the reboiler still are considered. The constant composition of distillate is considered to be same as in case of 1B, i.e., xD ¼ 0:572 . This is not much away from the equilibrium vapour composition with feed (xF ¼ 0.25, y ¼ 0:45). Four ideal stages were considered in 1B when the initial distillate concentration started from 0.8 and gradually fell; in this case only three stages are considered. The x-y diagram for McCabeeThiele construction is drawn for a ¼ 2.45. The initial operating line is constructed by trial and error with total three ideal stages starting from xB ¼ 0:25 and going up to xD ¼ 0:572 by construction of three stages as shown in Fig. 11.19B. From the y-intercept xD ( ¼ Rþ1 ¼ 0.336) of the line the corresponding reflux ratio required is calculated as 0.7016. As distillation proceeds, reflux ratio needs to be appropriately increased to keep xD ¼ 0:572, and xB decreases from 0.25. Therefore, several operating lines passing through (0.572, 0.572) are drawn and from the y-intercept their reflux ratio in each case is found out. xB is found in each case by constructing three stages on the operating line from (xD ; xD ). Only the initial and another operating line is shown in Fig. 11.19B that has y-intercept ¼ 0.151, and xB ¼ 0.16 read out from the graph. R ¼ 2.8 based on the intercept value. This is also considered to be the ending operating condition as the constant reflux ratio set for the previous case was also 2.8. ZxF Calculation of B=F : lnðB = FÞ ¼ xB

    dxB xD  xF 0:572  0:25 ¼ ln ¼ ln 0:572  0:16 ðxD  xB Þ xD  xB

ðB=FÞ ¼ 0:78 % recovery of toluene ¼ 0:78  0:16=0:25  100 ¼ 50 The reboiler total heat load may be found from the expression Z xB Q Rþ1 ¼ ðxD  xF Þ dxB 2 Flavg xF ðxD  xF Þ Observation on the results of this specific problem (1) Single-stage batch distillation is most inefficient with about 53% toluene recovery of just acceptable purity of 40% toluene in distillate. This would require minimum investment cost. (2) Batch distillation with constant distillate composition (w57% purity) is next best as it can be sold at a higher price. However, in this case the recovery is also low w50%. The distillation may be continued to further recover toluene but the incremental effect on yield increase would be lesser and lesser. The investment is more than single stage as two stages of contact in the column need to be provided. (3) Batch distillation with constant reflux ratio (w57% distillate purity) produces the same quality of product that can be sold at a higher price. However, in this case the recovery is much higher w83%.

350

Chapter 11 Distillation

For final decision a detailed economic analysis needs to be carried out that will include operating cost and the price differential between 40% purity toluene and >55% purity toluene may become a governing factor.

Further reading Nag, A. (2015). Distillation & hydrocarbon processing practices. PennWell Books. Kister, H. Z., Haas, J. R., Hart, D. R., & Gill, D. R. (1992). Distillation design (Vol. 1). New York: McGraw-Hill. Smith, B. D. (1963). Design of equilibrium stage processes. McGraw-Hill Companies. Treybal, R. E. (1980). Mass transfer operations. New York (p. 466). Brennan, K. (1990). The process engineers pocket handbook (Vol. 2). Houston, Texas: Gulf Publishing Company.

CHAPTER

Adsorption

12

12.1 Introduction Adsorption is one of the unique mass transfer processes in which component(s) from a fluid, either gas or liquid adheres to the surface of solid without intimate admixture with the solid atoms. As mentioned in Chapter 9, the solid on which adsorption occurs, often a porous material of a high specific surface (m2/g), is referred to as adsorbent while the substance that is adsorbed on the solid surface is termed as adsorbate or solute. All solids exhibit adsorption to some extent but certain substances exhibit preferential affinity towards specific solutes and can retain higher concentrations of those on their surface. This results in selective mass transfer and enrichment/separation/fractionation of components from a liquid solution or a gaseous mixture. Typical applications of liquid adsorption include adsorption of coloured matter from sugar solutions, petroleum, and vegetable oil; removal of objectionable taste and odour from potable water; removal of grease from dry cleaning liquids; removal of moisture dissolved in gasoline and fractionation of mixtures of paraffinic and aromatic hydrocarbons. Gas adsorption is widely used in industrial drying/dehumidifying air and other gases, removal of objectionable odours and impurities from industrial gases, recovery of valuable solvent vapour from dilute gaseous mixtures and fractionation of light hydrocarbons. In all these operations, the mixture to be separated is brought in contact with an insoluble solid, and the preferential distribution of the adsorbate on the (solid) adsorbent leads to the desired separation. Like other mass transfer operations, the contacting arrangement can be stagewise or continuous and the stagewise operation can be single-stage or multistage in crossflow or countercurrent mode. Since countercurrent operations are more efficient from the viewpoint of mass transfer, an approach to the same is achieved in some commercial adsorption processes that comprise an arrangement called simulated moving bed. MOLEX and PAREX are two such patented processes. Semi continuous mode of operation with the fluid flowing through a stationary bed of solid is also possible.

12.1.1 Modes of operation The modes of contact are presented as a classification tree in Fig. 12.1. For the purpose of calculation, the operation within each category except fixed-bed adsorption can be considered analogous to other mass transfer operations discussed in this book. For example, single-component adsorption from a gaseous mixture can be treated analogous to gas absorption where the added insoluble phase is adsorbent in the present case and liquid solvent in absorption. Single-component adsorption from

Process Equipment and Plant Design. https://doi.org/10.1016/B978-0-12-814885-3.00012-9 Copyright © 2020 Elsevier Inc. All rights reserved.

351

352

Chapter 12 Adsorption

liquid solution performed in batch, semicontinuous, or continuous mode can be treated analogous to mixer-settler operations of liquid extraction (contact filtration). Gas adsorption for solute removal or fractionation is usually performed in fluidised beds. The aforementioned analogies are often used to simplify the design procedure. Fixed-bed adsorption is an unsteady state rate controlled process, not similar to other mass transfer operations. Therefore, in this Chapter, we first outline the principles and governing equations for stagewise and continuous contact adsorption and then discuss in detail the fixed-bed adsorption process and its design.

Adsorption operation

Stagewise contacting

Single stage

Continuous contacting

Multistage

Cross-current flow

Fixed bed

Moving bed

Fluidised bed

Rotary bed

Counter current flow

FIGURE 12.1 Classification of adsorption operation based on the mode of contact.

Stagewise operation Single-stage adsorption for liquids is used for extremely favourable distribution of solute towards adsorbent. This is often employed to remove taste and odour from water using powdered activated carbon. For greater economy of adsorbent, the process is operated in multistage cross current or countercurrent mode Contact filtration accomplished by multiple tanks and filters connected in series. Countercurrent operation requires lower amount of adsorbent for the same separation, which is particularly important for expensive adsorbents like activated carbon. Although the savings are greater for larger number of stages, it is seldom economical to use more than two stages. A higher number of stages increase the expense of filtration and other handling costs. If the process needs to be operated in continuous mode as in decolorising petroleum lubricating oils, the filter press can be substituted by a centrifuge or a continuous rotary filter or the solid can be allowed to settle when the mixture is passed through a large vessel. Even if counterflow operation is more efficient and ensures greater adsorbent economy, cross current flow seems to be the more practical option in small-scale processing of liquids. This is particularly true when the amount of solution treated in every batch may vary considerably. Cross flow is also preferred when long intervals between batches may require partially spent adsorbents to be stored. Adsorbents like activated carbon

12.1 Introduction

353

deteriorate on longer storage due to oxidation, polymerisation and other chemical changes. The operation is usually performed at the highest convenient temperature to reduce viscosity and increase rate of diffusion and ease the motion of adsorbent particles through the liquid. The faster approach to equilibrium more than compensates the decrease in equilibrium concentration at the higher temperature. The highest permissible temperature can be close to the boiling point of the liquid provided the solid can withstand that temperature. Nevertheless, where the adsorption isotherm is strongly affected by temperature, the operation is best handled at room temperature. Temperature rise due to adsorption can be ignored when (i) the quantity of solution treated is much larger compared to the amount adsorbed, (ii) solute is adsorbed much more strongly compared to other constituents of solution and (iii) adsorbent is insoluble in solution. The component mass/mole balance can be expressed on solute free concentration basis, i.e., in terms of Y moles (mass) adsorbed /moles (mass) solute free solvent, and X moles adsorbed/mole solute free adsorbent. For single-stage operation, this yields a linear operating curve of slope ðAs =Ls Þ on the X Y plane, similar to the operating line for absorption. Ls ðYi  Yo Þ ¼ As ðXo  Xi Þ

(12.1)

In the above equation, Ls and As are the mass (moles) of solute free solvent and solute free adsorbent, respectively, and subscripts i and o represent the respective inlet and the outlet concentration. For an ideal stage, the adsorbate loading on the adsorbent is in equilibrium with the concentration of adsorbate in solution at the process temperature. Quite often equilibrium is given by Freundlich equation q ¼ KF ðpÞ1=Kc or q ¼ KF ðcÞ1=Kc , where p and c represent adsorbate partial pressure (in case of gas adsorption) and adsorbate concentration (in liquid adsorption) at exit. In this case, the amount of (fresh) adsorbent required to effect the change in feed concentration in a single ideal stage (equilibrium contacting) is given by As ðY0  Y1 Þ ¼ Ls ðY1 =KF Þ1=Kc For N stages in crosscurrent flow, the total amount of adsorbent required is " # As ðYN1  YN Þ 1=Kc ðY0  Y1 Þ ðY1  Y2 Þ ¼ KF þ þ :::: þ Ls ðY1 Þ1=Kc ðY2 Þ1=Kc ðYN Þ1=Kc

(12.2)

(12.3)

While for countercurrent operation with fresh adsorbent, the solute balance for N stages gives Ls ðY0  YN Þ ¼ As ðX1 Þ

(12.4)

and the adsorbent requirement in the N th stage is As ðYN1  YN Þ ¼ Ls ðYN =KF Þ1=Kc

(12.5)

354

Chapter 12 Adsorption

In Eqs. 12.2 e 12.5, the subscripts denote the exit concentration from the corresponding stage. For isothermal operation with fresh adsorbent, the calculation for theoretical stages can be made using Eq. 12.4. The same is also true for desorption/ regeneration. However, in practice the operation is adiabatic and the equilibrium relationship changes with stage number due to considerable rise (fall) in temperature during adsorption (desorption). Under this condition, the calculations are done stagewise similar to nonisothermal gas absorption or any stage-wise operation. For N stages of gasesolid contacting, solute balance in line with Eq. 12.4 is Gs ðY0  YN Þ ¼ As ðX1  XNþ1 Þ

(12.6a)

where Gs and As are the respective flow rates of solute free gas and adsorbent. The enthalpy balance gives   (12.6b) Gs ðHG0  HGN Þ ¼ As HS1  HS;Nþ1 where HG is the enthalpy of gas mixture in energy/mass   HG ¼ Cp;g ðtG  t0 Þ þ Y Cp;adsorbate ðtG  t0 Þ þ lvap;0

(12.7)

and HS is the enthalpy of the solid plus adsorbate in energy/mass adsorbent   HS ¼ Cp;adsorbent ðts  t0 Þ þ X Cp;adsorbent;liq ðtA  t0 Þ þ DHads

(12.8)

In Eqs. 12.8 and 12.9, enthalpies are expressed with respect to adsorbent, nonadsorbed gas and adsorbate as liquid, all at base temperature t0 , DHads is the integral heat of adsorption at X and t0 ; lvap;0 is the latent heat of vaporisation of adsorbate at t0 and Cp denotes specific heat of the component denoted by subscript.

Continuous contact operation Continuous contact devices can be operated in either continuous or semi continuous mode. In continuous mode, fluid and solid move continuously at a constant rate (steady-state moving bed adsorber) with either the solid moving relative to the walls of the container (moving beds) or the particles remaining stationary relative to the walls (rotary beds). In semi continuous mode, fluid moves through a stationary bed of solid and the composition changes with time (fixed bed adsorber). The merits and demerits of fixed and moving bed processes are listed in Table 12.1. In steady-state moving beds, the solid and the fluid phases move continuously through the equipment at a constant rate. At any point in the equipment, the composition of both the phases does not change with time. Use of moving bed is common for collecting solutes and fractionation of gaseous and liquid mixtures through Moving beds adsorption and ion exchange. Since the purpose of these applications is to achieve separation equivalent to many stages, only countercurrent operation is considered. Parallel flow may be used when one theoretical stage of contacting suffices. The major challenges of these equipment are (i) ensuring uniform flow of fluid and solid particles, (ii) uniform distribution of solid particles in the fluid to minimise channelling and local irregularities and (iii) continuous entry and removal of solids into the device. Attrition is unavoidable in moving bed processes and high mechanical (attrition) strength of adsorbent particles in the bed is a major requirement.

12.1 Introduction

355

Table 12.1 Merits and demerits of fixed and moving beds. Fixed bed

Moving bed

Advantages

· Simple inexpensive · Relatively · Minimal attrition of adsorbent particles · Easier to design

adsorbent inventory (regeneration · Low immediately after adsorption possible) · Better heat transfer

Disadvantages

· Difficult to design accurately due to the unsteady nature of · More complex and expensive equipment process · Attrition of adsorbent inventory of expensive adsorbent and · Large proportionately large pressure drop for the slow progress · ·

a

of MTZa Safety concern for highly porous adsorbents as poor thermal conductivity may lead to hot spots and fire hazards, e.g., in case of activated carbon bed with air. Also, the equilibrium may be adversely affected. Multiple beds required for continuous operation.

Mass transfer zone (MTZ), elaborated later in the chapter.

For single component adsorption, the computations can be performed using Eq. 12.1 Gs ðYi  Yo Þ ¼ As ðXo  Xi Þ   which gives a linear operating line of slope As =Gs . The line lies below the equilibrium curve for adsorption and above it for desorption. Minimum adsorbent requirement is given by the operating line of maximum slope that touches the equilibrium curve at a point as discussed in Chapter 10 (Section 10.3). Although in practice, the operation is adiabatic and the temperature does not remain constant except for solute collection from vey dilute liquid solutions, the calculations are usually done assuming isothermal operation following the design procedure of packed bed discussed in Chapter 10. The governing equations for gas adsorption are As dX ¼ Gs dY ¼ KY aP ðY  Y  Þdz

(12.9)

Which on rearranging gives ZY1 NtOG ¼ Y0

dY KY aP ¼ Y  Y Gs

Zhbed dz 0

(12.10a)

356

Chapter 12 Adsorption

where HtOG ¼

Gs KY aP

(12.10b)

Y  is the equilibrium composition in the gas corresponding to the adsorbate composition X and KY aP is the overall gas-phase mass transfer coefficient based on the specific surface area of the solid particles ðaP Þ. Rotary beds are employed to combine the advantages of both moving and fixed bed. Several beds are on a rotating drum (or wheel) that makes each bed to go through cycles of adsorption and desorption. At specific locations, each bed aligns with the appropriate inlet and outlet piping connections for adsorption and desorption. Rotary beds Most often these beds are partitioned sectors. Recovery of volatile solvent from air is one of the common applications of rotary bed adsorber. Fluidised beds are extensively used for recovery of vapours, drying of air with silica gel, fractionation of light hydrocarbon vapours with carbon, etc. Fluidised beds are attractive because (i) interparticle heat and mass transfer is high and (ii) when fully fluidised, its net weight is supported by drag forces due to fluid flow that Fluidized beds results in the pressure drop (DP) to be almost independent of flow rate (Eq. 12.11). DP ¼ ð1  aÞðrs  rf Þg Lbed

(12.11)

where rs and rf are the density of the solid and fluid, respectively, and a is the bed voidage at the onset of fluidisation.

12.1.2 Adsorption mechanisms Bonding of a solute on a solid surface can be physical bonding (Physisorption) or chemical bonding (Chemisorption). Table 12.2 lists the major differences between the two processes. There may be situations when bonds of both types may be present. A third mechanism e capillary condensation, happens in the case of gas adsorption in porous media. This involves localised condensation inside pores at temperatures above the dew point of the bulk fluid allowing the formation of multiple solute layers on the surface. Porous adsorbents with capillaries not too narrow on the molecular scale adsorb by the same mechanism as nonporous adsorbents for a low relative partial pressure (partial pressure/vapour pressure). With increasing pressure, multilayer adsorption takes place with adsorbate condensing Adsorption in Porous media inside pores and the heat of adsorption similar to the heat of condensation. This phenomenon termed “capillary condensation” can be explained by considering the vapour pressure over a curved surface. At the same temperature, the vapour pressure is lower over a concave surface as compared to a flat or convex surface. In the narrow capillaries, the adsorbed liquid meniscus is concave. This results in the vapour condensing at a lower pressure than the vapour pressure over a flat surface. Capillary condensation causes hysteresis of adsorption equilibrium (Fig. 12.2) in porous adsorbents.

12.1 Introduction

357

Table 12.2 Physisorption versus Chemisorption. Physisorption

Chemisorption High enthalpy of adsorption (200e400 kJ/mol)

Reversible

Irreversible

Weak forces of attraction like van der Waals forces, hydrogen bonding, etc.

Chemical bonding involving orbital overlap and charge transfer

Multilayer adsorption. BET isotherm used to model equilibrium.

Generally, monolayer adsorption Langmuir isotherm used to model equilibrium

Observed at low temperature (higher temperature reduces surface coverage)

Observed at higher temperature

Not specific No surface reactions

Highly specific adsorbate-adsorbent pairs Surface reactions e dissociation, reconstruction, catalysis possible

p*, Equilibrium partial pressure

Low enthalpy of adsorption (5e50 kJ/mol)

Adsorption

Desorption

q, kg adsorbate/kg adsorbent

FIGURE 12.2 Adsorption isotherm exhibiting hysteresis.

12.1.3 Adsorption equilibrium Isotherms relating adsorbate loading capacity to adsorbate concentration in the process stream at a given temperature are the most common form of expressing the equilibrium data. Equilibrium data governs capital cost to a large extent since adsorption capacity decides the amount of adsorbent required.

358

Chapter 12 Adsorption

Adsorption isotherms are influenced by: • • • •

Nature of adsorbent Cycles of adsorption and desorption which alter adsorbent characteristics possibly due to progressive changes in the pore structure Origin and method of preparation of adsorbent Temperature and relative humidity of the vapour/gas stream

Diffusive characteristics in liquids significantly affect the adsorbent performance, and vapourphase isotherms are more readily available than liquid-phase applications. The designer, however, needs to judiciously consider the equilibrium relationship in the gas phase since it is affected by hysteresis as well as loss in adsorption capacity with cycles of regeneration of the adsorbent. It is, therefore, a common practice to use adsorption isotherms developed from experiments on a pilot scale before designing industrial adsorbers. The phenomena of hysteresis observed in gas adsorption is illustrated in Fig. 12.2 which shows that the equilibrium path followed during adsorption and desorption (from the final state of adsorption) are different. The desorption (equilibrium) pressure is always lower than the corresponding adsorption pressure. The phenomenon arises when Hysteresis adsorption occurs primarily following the capillary condensation mechanism and can be attributed to the difference in liquid meniscus shape (curvature) during adsorption and desorption. Spherical and cylindrical menisci are formed during adsorption whereas during desorption the menisci are spherical. Hence, desorption isotherm is used to determine effective pore size. Due to the absence of capillary condensation during liquid-solid adsorption, liquid adsorption is usually reversible and does not exhibit hysteresis. However, there may be a loss in capacity with cycles of regeneration that the designer has to consider.

12.2 Packed bed adsorption Packed beds are widely used with both gas and liquid feeds. Separation in a fixed bed is an unsteady state rate-controlled process and at a particular axial location within the bed, the conditions vary with time. At the start of the process, as feed enters the bed, mass transfer occurs near the inlet and concentration of the adsorbent in the fluid phase decreases from inlet value to near-equilibrium concentration over a narrow zone. This portion of the bed is termed the mass transfer zone. With progress of time, the initial part of the mass transfer zone (MTZ) becomes almost saturated and is unable to adsorb further solute (equilibrium zone). The “unadsorbed” adsorbate then gets carried further downstream, and thus, the mass transfer front proceeds in the direction of feed flow while the rear end of the mass transfer zone gets saturated. The net effect is a forward movement of MTZ, leaving behind an equilibrium zone saturated with solute. The portion of the bed beyond MTZ is not yet in contact with the solute and is, therefore, unutilised. It is capable of mass transfer and is termed the active zone. Thus, at any instant of time, the entire bed can be divided into three zones based on mass transferdequilibrium zone, mass transfer zone, and active zone, as illustrated in Fig. 12.3A. At any instant, adsorption is confined to MTZ, and the adsorbent upstream or downstream of MTZ do not participate in the adsorption process. As the fluid continues to flow, the mass transfer zone (MTZ) moves downward as a wave (Fig. 12.3B) at a rate usually much slower than the fluid velocity and the effluent concentration is substantially zero, till MTZ reaches the effluent end of the bed.

12.2 Packed bed adsorption

(A)

Mass Equilibrium transfer zone zone

Feed Ci

(B)

Active zone

359

Effluent

1 t1

C/Ci 0

t2

C* Ci

t3

t4

Distance along bed (z) Concentration in fluid phase are shown at time instants t1, t2, t3, t4

FIGURE 12.3 (A) Zones in packed bed during adsorption (at time instant t2); (B) Solute concentration profiles in the fluid phase with progress of time (c is equilibrium concentration).

12.2.1 Breakthrough curve, breakthrough point, and bed exhaustion As adsorption in a packed bed continues, adsorbate concentration in bed effluent remains “zero” as long as the MTZ remains within the bed and has at least an infinitesimally small active bed in front of it. Beyond this point of time, as the MTZ moves further downstream and exits the bed, adsorbate starts slipping out Breakthrough curve of the bed along with the effluent. This is observed as an increase in the concentration of adsorbate in the effluent stream with time from “nil” or a “negligible value.” The, typically “S-shaped,” concentration versus time curve shown in Fig. 12.4 is called the concentration breakthrough curve. The heat released during vapour adsorption from a gas mixture is not quickly dissipated due to the lower thermal conductivity of the gas phase and also the porosity of the adsorbent. This increases the temperature locally, and a temperature wave similar to the adsorption wave is generated. The rise in

C/Ci

1.0 C/Cex

C/Cb 0

Bed exhaustion

Bed saturation

Breakthrough point Time, t

FIGURE 12.4 Concentration breakthrough curve.

360

Chapter 12 Adsorption

temperature at the fluid outlet can thus be used as a rough indication of breakthrough point. The temperature rise is relatively small for liquid adsorption as the heat gets more easily dissipated. The onset of presence of adsorbate in the effluent is the ideal “breakthrough point.” However, in practice, the effluent stream is inevitably used in a downstream process and the maximum allowable solute concentration in the stream acceptable by the said process is the “breakthrough concentration” considered in bed Breakthrough point design. Operation beyond the breakthrough concentration renders the effluent stream unacceptable for the downstream process. Operation beyond the breakthrough point continuously increases the concentration in the effluent stream, and finally, the entire bed gets fully “exhausted.” No further adsorption takes place in this condition, and the effluent stream concentration is close to the feed stream concentration. In practice, the bed is considered Bed exhaustion exhausted when the effluent concentration is around 95% of the feed concentration. One may also note that when some very strongly adsorbed components are present in feed stream along with a mixture of less strongly adsorbed components, the effluent concentration may not attain the feed concentration as only the components that move fast through the bed, are in the breakthrough curve. Shape of the breakthrough curve strongly depends on (i) Rate and mechanism of the adsorption process (ii) Nature of adsorption equilibrium (a flat isotherm indicates a narrow MTZ and a steep breakthrough curve) (iii) Fluid velocity; Concentration of adsorbate in feed (iv) Mass transfer kinetics: fast kinetics implies a steep breakthrough curve while slow kinetics results in a distended shape. Slow kinetics may be tackled by increasing cycle time and also by using smaller adsorbent particles, but the designer needs to consider that increased cycle time requires a higher inventory of adsorbent and smaller particles entail higher pressure drop. Breakthrough capacity (BC) of a bed is the mass of adsorbate held per unit mass of adsorbent at the point of breakthrough, and this clearly depends on the breakthrough concentration (cb,) at breakthrough time (tb) (Fig. 12.4) considered by Breakthrough capacity and Saturation capacity the designer. Similarly, the saturation capacity of the bed (SC) is the maximum loading of adsorbate when the entire bed is fully saturated and is known from the equilibrium data/exhaustion point (cex). The monotonically increasing concentration profile in the mass transfer zone is considered as symmetric around its midpoint, and hence, the average bed loading in the mass transfer zone of length MTZ is 50% of SC. This leads to the following relationship for a bed of length L   ½L  MTZ MTZ þ 0:5 BC ¼ SC (12.12) L L Industrial beds go through cycles of adsorption and regeneration, and some residual solute loading (RC) remains in the bed after regeneration. Designers incorporate the effect of RC and use working

12.2 Packed bed adsorption

361

capacity WC as a measure of the actual adsorption capacity of the bed. If experimental data is available, WC may be estimated as   ½L  MTZ MTZ þ 0:5 WC ¼ SC  RC (12.13) L L In the absence of data WC is taken as a fraction of SC as WC ¼ SC  f

(12.14)

Typical value of f may be 0.85e0.9. During operation, the adsorption step is terminated slightly before the breakthrough point. This ensures that the effluent always remains “on-spec.” In planning new processes, it is best to determine the breakthrough point and breakthrough curve for a particular system experimentally under conditions as close as possible to the process conditions.

12.2.2 Desorption/regeneration The saturated adsorbent is either regenerated or disposed off. Disposal may be considered as an option when (1) adsorbent cost is low, (2) regeneration is very difficult/expensive, (3) nonadsorbed component is a very high value desired product, (4) chemisorption occurs and reversibility is impractical. In most applications, disposal is uneconomic, and adsorbent is regenerated for reuse either in-situ or in a separate process. The environmental effect of the disposed adsorbent is also a concern for the designer.

Gas-phase adsorption

q1

Adsorbent loading

Adsorbent loading

Regeneration of gas adsorbers, i.e., the desorption step, involves changes in temperature, adsorbate partial pressure, or passing a competitively adsorbing component through the bed. The effect of temperature and pressure on equilibrium loading can be understood from Fig. 12.5. Reducing partial pressure of solute from p1 to p2 reduces equilibrium loading from q1 to q2 (Fig. 12.5A). At constant partial pressure/concentration of the adsorbate in the gas phase (or concentration in the liquid phase), an increase in temperature from T1 to T2 decreases the equilibrium loading from q1 to q2 in Fig. 12.6B.

q2

p2

p1 Partial pressure

q1

T1

q2

T2

q3

T2>T1

p2

p1

Partial pressure

FIGURE 12.5 Effect of process variables on adsorption equilibrium for a Type I isotherm: (A) Adsorbate partial pressure (PSA pathway) (B) Temperature (TSA pathway).

362

Chapter 12 Adsorption

Further reduction of partial pressure from p1 to p2 at temperature T2 reduces the concentration to an even lower value of q3. In practice such combinations of depressurising and heating is employed for desorption. The final choice of the regeneration method is based on technical and economic considerations (capital and operating cost). A brief description of the processes is provided below, and the general comparison of the methods along with specific applications is listed in Table 12.3. In PSA, the partial pressure of adsorbate may be reduced by reducing the total pressure (evacuating the bed), and also by introducing an inert component (diluent) through the bed while keeping the total pressure same. In many practical processes, both are used in Pressure Swing Adsorption (PSA) sequence. If the reduction in pressure involves a vacuum, the process is often called Vacuum Swing Adsorption (VSA). The use of only a diluent stream for regeneration is not very popular. Since pressure change is a quick process, the cycle time of PSA systems is low, usually in minutes. One may also note that PSA (and VSA) can be controlled either by equilibrium or by kinetics depending on the adsorbent. Both types are important commercially, and a typical example is air separation. With zeolite as adsorbent, nitrogen is adsorbed more strongly than oxygen, and the equilibrium controlled separation produces oxygen of almost 96% purity as product. On the other hand, when using carbon molecular sieves, the equilibrium loading of both nitrogen and oxygen are close but very high purity (99%) nitrogen is obtained in the effluent, as oxygen, due to its high diffusivity, gets rapidly adsorbed. Regeneration in the TSA process is achieved by increasing the temperature at a constant partial pressure in the gas phase or concentration in the liquid phase. Since Temperature Swing Adsorption (TSA) the strongly adsorbed components have a high heat of adsorption, an increase in temperature causes a substantial decrease in their loading. Hot purge gas or steam is almost always used along with bed heating. The use of purge gas flushes the desorbed component(s) from the bed and also reduces its partial pressure that aids further desorption. Thermal processes are always slower, and so TSA processes operate with longer cycle time. This is a major reason for using TSA virtually exclusively for treating feeds with a low concentration of adsorbate. Temperature cycling of TSA beds lead to cycles of expansion and contraction that creates stress on the adsorbent particles and generation of fines that may increase bed pressure drop. Also, accidental running with high temperature may cause deactivation of bed. Therefore, degradation of adsorbent is a cause of concern in TSA. Some adsorbates are difficult to separate by PSA or TSA. These can be desorbed by displacement with a more preferentially adsorbed species that can be a gas, Displacement Purge Adsorption (DPA) vapour or liquid. The choice of the displacement fluid is important as a displacement fluid adsorbed too strongly may be difficult to remove from the adsorbent in the next step. Desorption occurs due to a reduction in partial pressure (or concentration) of original adsorbate in the fluid phase, as well as competitive adsorption of the displacing component. Since adsorption and desorption go on simultaneously, the net heat effect in the bed is small, and this keeps a more or less constant bed temperature throughout the

12.2 Packed bed adsorption

363

Table 12.3 Comparison of the typical regeneration processes.

Method

Salient characteristics

Typical processes

Typical adsorbent & process

Advantages

Disadvantages

TSA

Suitable for both gases and liquids. Good for strongly adsorbed species. Desorbate recovered at high concentrations.

Thermal aging of adsorbent. Adsorbent degradation at very high temperatures. Long cycle times (several hours). High initial cost due to heating arrangement. Consumes large energy per unit adsorbent quantity. More complex control need skilled personnel for maintenance.

Always used in conjunction with hot gas/steam purge. Used virtually exclusively for feeds with low adsorbate concentration.

Drying of gases/ organic vapour. Drying of solvents.

Molecular sieve; activated carbon

PSA

Good for weakly adsorbed species required in high purity. Lower adsorbent inventory and smaller size of adsorbent bed lower cycle time

Cannot be used for liquids. Very low pressure may be required. Mechanical energy more expensive than heat

Operates close to ambient temperature.a Low adsorbent loading.b

Drying of gases; Hydrogen recovery; Bulk gas separation (air separation).

Molecular sieve; carbon molecular sieve; zeolite.

VSA

Rapid cycling gives efficient use of adsorbent.

Desorbate recovery at low purity.

Separation of linear paraffins.

Molecular sieve.

Displacement

Isothermal operation. Good for strongly held species. Avoids risk of cracking reactions during regeneration Avoids thermal aging of adsorbent.

Product contamination by displacement fluid. Product separation and recovery necessary.

Separation of linear from cyclic and branched paraffins.

Molecular sieve

Adsorption of displacement fluid almost as strong as desorption of adsorbate.

Continued

364

Chapter 12 Adsorption

Table 12.3 Comparison of the typical regeneration processes.dcont’d

Method

Advantages

Disadvantages

Purge gas stripping

Essentially at constant T and P. Simple process. Little maintenance. Cheap (low capital cost) Safe.

Only for weakly adsorbed species. High purge flow. Usually not used when desorbate needs to be recovered. High operating cost in large systems.

Steam stripping (combination of TSA and displacement)

Same as TSA and displacement

Same as TSA and displacement

Salient characteristics For small capacity systems (smaller than 3500 m3/hr approx.).

Typical processes

Typical adsorbent & process

Relatively uncommon without thermal swing. Purging alone suitable for only weakly adsorbed species.

Waste water purification; Solvent recovery.

Activated carbon.

a Typically PSA processes are operated close to ambient temperature since the loading increases as temperature is decreased at a constant partial pressure. b PSA processes are often operated at low adsorbent loading since selectivity among gaseous components is often greatest in Henry’s law region.

cycle. The displacement purge also flushes the desorbed components out of the bed. The contamination of the products by the displacement fluid is a disadvantage of the process. The cost of further separation of the contaminated product, if involved, needs to be considered in assessing the overall economics of the process.

Liquid-phase adsorption PSA is inapplicable in this case, as the effect of pressure on equilibrium is minimal. Since the equilibrium in the case of liquid-phase adsorption is sensitive to temperature, TSA is used for desorption in some cases. Hot gas is often used for heating. In these cases, arrangements for bed draining after the adsorption step and refilling it with liquid are provided. Table 12.4 lists process features and the feasible regeneration techniques for each. For an assigned adsorption operation, one needs to identify the relevant features pertaining to the operation and then note the techniques denoted by Y/N for each. The techniques with N are eliminated, and a process with Y for all conditions is the best choice. Entries other than Y/N are used to rank processes if more than one option seems to be feasible for the said operation.

12.2.3 Adsorbent aging Adsorbent capacity gradually reduces during the working life of adsorbent. It occurs due to multiple regeneration and results from loss of active surface area. In commercial units, hydrothermal

Table 12.4 Cyclic adsorption process options for various conditions.

Sl. No.

TSA b

c

G

Process conditions

L

G

L

Inert purge

SMBa

Chromatography

Displacement

Only for L

PSA

Only for G

1

Liquid feed that can be completely vaporised below 200 C

Not likely

Y

Y

Y

Y

Y

Y

Y

2

Liquid feed that cannot be fully vaporised below 200 C

N

Y

N

Y

Y

N

N

N

3

Adsorbate concentration in feed 10 wt%

N

N

Y

Y

Y

Y

Y

Y

6

Adsorbate recovery at high purity->90%e99% rejection of carrier

Y

Y

Y

Y

Y

Y

Y

Maybed

7

Adsorbate only desorbed by TSA

Y

Y

N

N

N

N

N

N

8

Practical (cheap, noncorrosive nontoxic adsorbate)

Y

Y

N

N

N

Y

Y

Y

9

Displacement or purge agents not easily separated from adsorbate

May be

May be

Not likely

Not likely

Not likely

Not likely

Not likely

N

10

Vaporised liquid/Gaseous feed

Y

N

Y

N

N

Y

Y

Y

a

SMB e Simulated Moving Bed. G e Gas-Phase Applications. c L eLiquid-Phase Applications. d Greater than 10:1 ratio of feed to desorption pressure/vacuum desorption required. b

366

Chapter 12 Adsorption

deterioration and chemical contamination/fouling are the main causes. The hydrothermal effect due to water exposure at regeneration temperature results in “collapsed pore structure” and is irreversible. Contamination occurs when the adsorbent active surface is blocked by direct deposit or by degradation, polymerization, or oxidation of unstable compounds present in the fluid. Although this form of aging is considered theoretically reversible, in practice, the immobile adsorbent deposits increase at each regeneration and decrease adsorption capacity. Formation of fines with time, particularly after each regeneration, also increases bed pressure drop with time. Longer intervals between regenerations, therefore, allow greater utilization of the available capacity. This may be done by operating columns in series such that the first column is operated until it is fully saturated and the second column ensures that the specification is met.

12.2.4 Bed design Fixed-bed adsorbers can be designed either by the rigorous solution of conservation, transport, and thermodynamic equations or by experimental breakthrough curve data generated in the laboratory scale, pilot scale, or industrial scale, and scaling up for the design of the actual column. Data required for design by rigorous methods is often unavailable, and predictive modeling of heat, mass, and momentum transfer are resorted to. Accuracy of prediction depends on the availability and reliability of fundamental data and the validity of the assumptions and approximations to obtain the solution. These solutions are usually obtained by numerical analysis.

Rigorous methods Due to the unsteady nature of the adsorption process, the mathematical model is in the form of partial differential equations. Both energy balance and mass balance equations are to be considered for an accurate design. However, many fixed-bed designs consider an isothermal operation that is valid for low concentration of adsorbate and low heat of adsorption. This allows the design to be based on mass balance only. The fluid phase unsteady-state mass balance for a single adsorbate (from liquid feedstock) is   v2 c vðucÞ vc 1  a vq  DL 2 þ þ þ rp ¼0 (12.15) vz vz vt a vt where the first, second, third, and fourth terms represent axial dispersion (DL is axial dispersion coefficient), convective flow within the bed (u is the interstitial velocity, i.e., the superficial velocity/ void fraction a), accumulation of adsorbate in the fluid phase and rate of adsorption, which is a function of fluid phase concentration and adsorbent loading. vq ¼ f ðq; cÞ (12.16) vt The equation is used for gas-phase adsorption by replacing the concentration terms with adsorbate partial pressure. In that case, a pressure drop equation to estimate total pressure along the bed may be necessary as it will affect the adsorbate partial pressure. If the bed pressure drop is small compared to the total pressure, the average bed pressure is considered. Simultaneous solution of Eqs. (12.15) and (12.16) along with the initial and boundary conditions generate the transient response of the bed. The shape of the MTZ depends on the shape of the

12.2 Packed bed adsorption

367

equilibrium isotherm, concentration of an absorbable component, hydrodynamics, and choice of a kinetic model. The simplest design case considers a single dilute adsorbate in an isothermal process for axially convected plug flow and no mass transfer resistance. This simplifies Eq. (12.15) to   vc vc 1  a vq ¼0 (12.17) u þ þ rp vz vt a vt And the equilibrium isotherm is given by q  ¼ f ðcÞ

(12.18)

An a priori design of cyclic TSA and PSA processes is not very reliable as the mathematical models of all the steps are not sufficiently detailed to capture the process fully. Much of the input data necessary are also estimates and mostly tested on small-scale laboratory set up. Hence, industrial adsorbers are often designed by scale up of pilot plant data using empirical methods.

Empirical or short-cut methods The usual practice is to design adsorbers based on bench scale isotherm data and pilot plant data on breakthrough characteristics where the pilot operating characteristics (adsorbate loading rate, detention time, superficial velocity through the bed) are as close as possible to those expected in the scaled up system. The data on breakthrough characteristics (breakthrough curve and time of breakthrough, % approach to saturation at breakthrough) corresponding to selected process parameters like surface loading rate and empty bed contact time generate scale up factors for full-scale design. These are utilized along with the following assumptions: (i) (ii) (iii) (iv) (v) (vi)

isothermal adsorption from dilute feed mixtures adsorption isotherm is concave to solution concentration axis constant length of MTZ as it travels through the adsorber bed height of adsorber bed is large compared to the depth of MTZ axially dispersed plug flow negligible mass transfer resistance

Many industrial adsorbers conform to these assumptions.

Pilot plant design Many of the practices valid for the scaled up version are valid for the pilot plant as the scale up is based on similarities in mass transfer. Pilot plant studies are almost always desirable for liquid-phase applications due to the high sensitivity of site-specific parameters like change in pH, etc. An accurate scale up requires the following considerations: - Same fluid and adsorbent pair and adsorbent particle size, shape porosity, and other characteristics. Similar hydrodynamics and dispersion characteristics as actual bed. Mal-distribution/channelling, usually more pronounced at lower flow rates should not occur. Designers attempt to minimise channelling by keeping the ratio of bed diameter (D) to particle diameter (dp) above a minimum limit. This limit varies from 20 to 30. Also, the minimum ratio of adsorbent bed length (L) to the

368

-

-

Chapter 12 Adsorption

adsorbent particle diameter (dp) is kept at least 100. This is done to allow the mixing of liquid at the wall and the rest of the bed. To summarise, the designer respects (D/dp) > 20 to 30, and (L/dp) > 100. More than one MTZ will occur for all isotherms e favourable, unfavourable or mixed for multicomponent and/or nonadiabatic systems and it is a conservative approach to consider the same bed length for small-scale and scaled up process. Length of the pilot-scale bed needs to cover several mass transfer zone lengths.

It may be noted that breakthrough curves are different for adsorption and desorption (Fig. 12.6). The designer usually considers the breakthrough curve pertaining to adsorption and verifies the adequacy of design for the desorption step. Identical filtration rate (FR) and empty bed contact time (EBCT) for full scale and pilot plant ensure similar mass transfer characteristics and superficial velocity (US) defined as volumetric flow rate/cross-sectional area of the empty bed. For scale up with the same superficial velocity, the cross-sectional area of the actual plant needs to be increased. This can be either by increasing the number of beds operating in parallel and/or using a single bed with a large diameter. The first option requires higher capital cost but ensures identical operation of the full-scale and small-scale unit. On the other hand, the second is a cheaper option but requires proper flow distribution and redistribution arrangements.

Adsorption

1

2

3

C/Ci

1: Minimal diffusion 2: Moderate diffusion 3: Moderate dispersion and diffusion

3

2 1 Time

2

Desorption (Regeneration) 1

C/Cinit

3

1: Minimal diffusion 2: Moderate diffusion 3: Moderate dispersion and diffusion

3 1

2

Time

FIGURE 12.6 Typical breakthrough curves during adsorption and desorption (ci is the concentration of adsorbate in feed and cinit is the concentration at the start of the desorption cycle).

12.2 Packed bed adsorption

369

Data/information required for design - Choice of adsorbent; Equilibrium data, appropriate interaction data for multicomponent systems/ effect of other adsorbates on equilibrium; Selection of desorption method - Adsorbent-adsorbate kinetics/suitable interaction data for multicomponent systems - Heat of adsorption at the operating conditions. This may be used to check if the isothermal operation can be assumed for design - Hydrodynamic data to estimate pressure gradient - Property data over the operating range Typical commercial adsorbents and their applications are listed in Table 12.5. Properties of some common adsorbents are listed in Table 12.6. Usually, aqueous solutions are treated with freshly prepared activated carbon and organic liquids (oils) are adsorbed by clays (inorganic adsorbent). OccaCommercial adsorbents sionally mixed adsorbents are used.

Table 12.5 Commercial adsorbents and their typical applications. Adsorbent

Application

Zeolite

Separation of normal paraffin (adsorbate)/isoparaffin/aromatics; N2 (adsorbate)/O2; H2O (adsorbate)/ethanol Adsorption of CO2 from C2H4 natural gas, etc; SO2 from vent streams; sulphur compounds from organics, natural gas, H2, LPG, etc.; NOx from N2; water from olefin containing cracked gas, natural gas, air, syn gas, etc. Liquid bulk separation of n-paraffin/isoparaffin/aromatics/p-xylene (adsorbate) from o-xylene, m-xylene; detergent-range olefins (adsorbate) from paraffins; (adsorbate); fructose (adsorbate) from glucose; sulphur compounds (adsorbate) from organics

Activated carbon

Gas-phase separations of ethylene and organics (adsorbate) from vent streams; solvent/odours (adsorbate) from air stream; Removal of organics, oxygenated organics, chlorinated organics, etc. from water; odour, taste bodies from drinking water; fermentation products from fermenter effluent; decolorising petroleum fractions, sugar syrup, vegetable oil, etc.

Alumina

Water removal from organics, oxygenated organics, chlorinated organics, olefin containing cracked gas, natural gas, air syn gas, etc.

Silica

Water removal from organics, oxygenated organics, chlorinated organics, olefin containing cracked gas, natural gas, air, syn gas, etc.

Carbon Molecular sieve

Separation of O2 (adsorbate) from N2

370

Chapter 12 Adsorption

Table 12.6 Physical properties of typical commercial adsorbents. (a) Adsorbent grade activated alumina Bulk density of 800 kg/m3, pore size 1e7.5 nm, pore volume 0.40 cc/g Adsorption properties

% w/w

H2O capacity at 4.6 mm Hg, 25 C

7

H2O capacity at 17.5 mm Hg, 25 C

16

25 C

CO2 capacity at 250 mm Hg,

2

(b) Adsorbent grade silica gel Bulk density of 250 kg/m3, pore size 1e7.5 nm. Pore volume 0.40 cc/g Adsorption properties

% w/w

H2O capacity at 4.6 mm Hg,

25 C

11

H2O capacity at 17.5 mm Hg, 25 C CO2 capacity at 250 mm Hg, O2 capacity at 100 mm Hg,

35

25 C

3

(183 C)

22

25 C

17

n-C4 capacity at 250 mm Hg, (c) Activated carbon Physical properties

Liquid-phase application

Vapour-phase application

Wood based

Coal based

Wood based

Coal based

Bulk density (kg/m )

250

500

500

530

Mesh size (Tyler)

100

8þ30

4þ10

6þ14

Ash (%)

7

8

8

4

3

Adsorption properties H2O capacity at 4.6 mm Hg,

% w/w 25 C

11

H2O capacity at 250 mm Hg, 25 C n-C4 capacity at 250 mm Hg,

5e7

25 C

25

(d) Molecular sieve Zeolite type Designation Pore size (nm) 3

Bulk density (kg/m )

A LiA/KA

NaA

CaA

X NaX

3A

4A

5A

13X

0.3

0.4

0.5

0.8

670e740

660e720

670e720

610e710

12.2 Packed bed adsorption

371

Table 12.6 Physical properties of typical commercial adsorbents.dcont’d O2 capacity at 100 mm Hg, (183 C)

Not adsorbed

22 %w/w

22 %w/w

24 %w/w

H2O capacitya at 4.6 mm Hg, 25 C

20 %w/w

23 %w/w

21 %w/w

25 %w/w

CO2 capacitya at 250 mm Hg, 25 C

Not adsorbed

13 %w/w

15 %w/w

16 %w/w

n-C4 capacitya at 250 mm Hg, 25 C

Not adsorbed

Not adsorbed

10

12

a

% wt on activated pellet.

Particle sizes for adsorbents typically range from 100 or 200 mesh, going up to 6.5 mm nominal size. The typical size and shape factors of adsorbents are given in Table 12.7. Bead size is also denoted by screen analysis as it is difficult to manufacture beads of a uniform Choice of Particle Shape and Size size. Higher mass transfer rates are obtained for high specific surface area and lower pressure drop results for larger shape factor. Table 12.7 Adsorbent size and shape factors. Particle shape

Granules

Beads

Pellets

Typical size

100/200 to 4/8 mesh screen analysis 50 mm to 12 mm diameter

16/40e4/8

0. 4e0. 8 mm

Shape factor (j)

0.45e0.65

1

0.63

(DP/L)N jn3; n ¼ 1 for laminar and 2 for turbulent flow

For any application, particle size is selected from the range of commercial adsorbent sizes available and its effect on (1) mass transfer, (2) pressure drop, (3) axial dispersion. While larger particle size is favoured from pressure drop and axial considerations, they result in a lower rate of mass h dispersion  i transfer leading to larger bed size ðkc aÞN dp1.5 to 2.0

. Thus, the optimum size is an economic

balance between operating costs due to pressure drop and capital investment. Typically, 4/6 or 4/10 mesh carbon is selected for gas purification/adsorption, and pressure drop is not a problem. Smaller particles are chosen for liquid adsorption, e.g., 20/50 mesh carbon (0.3 mm  dp  .85 mm) is chosen for water treatment. In some applications, two sizes of adsorbents are used to trade-off between pressure drop and mass transfer rate. Larger sized particles are placed at bed inlet since mass transfer effects are more important towards the effluent end after MTZ terminates.

372

Chapter 12 Adsorption

Spherical particles are preferred over pellets and extrudates as they enable higher flow rates and taller towers due to (i) Lower pressure drop per unit volume (ii) Comparatively uniform and compact loading with little subsequent settling (iii) Lower attrition and crushing Few specific gas adsorption applications use the adsorbent slurried in liquid for ease of handling and sparged vessels used for contacting may operate in semibatch mode (continuous gas flow), concurrent flow, or both. A typical application is adsorption of SO2 from the air-SO2 mixture using activated carbon slurry in water. It has been shown that the capacity of adsorbent remains unchanged in dry or in slurry and is much larger than that of liquid solvent alone.

Operating parameters from pilot tests (a) Loading rate/filtration rate (LR) for liquid-phase applications Loading rates are usually 80e240 lpm/m2 of bed cross-section. Occasionally high rate up to 400 lpm/m2 may be used. Bed diameter D is sffiffiffiffiffiffiffiffiffiffiffiffiffi 4Q D¼ (12.19) pðLRÞ Where Q is the fluid volumetric flow rate through the bed and LR ¼ LRpilot ¼ ðQ=AÞpilot

(12.20a)

One needs to note that the variables have to be in consistent units and subscript pilot refers to pilot plant data.

(b) Superficial velocity (Us) for gas-phase applications The diameter of the adsorber vessel decides the superficial velocity, i.e., the volumetric inflow rate/ cross-sectional area of the empty bed. This influences the MTZ by affecting the kinetics of solute transfer from the fluid to the adsorbent surface. ðUs Þ ¼ ðUs Þpilot ¼ ðQ=AÞpilot

(12.20b)

Beds are typically designed with Us between 0.25 and 0.6 m/s. Using a higher velocity results in a lower diameter. This is often not desirable, the reason is not entailing higher pumping cost but lower residence time. In case of pressure drop limitation, Us can be lowered up to 0.12 m/s.

(c) Empty bed contact time Empty bed contact time (EBCT) is related to the contactor dimension as       V LA LA EBCT ¼ ¼ ¼ Q Q Q pilot

(12.21)

and provides an estimate of actual bed depth from pilot plant data where V, L, and A are the volume, length, and cross-sectional area of the adsorbent bed while the corresponding nomenclatures

12.2 Packed bed adsorption

373

subscripted as “pilot” refer to the pilot plant data. Contact time specified for drying liquids are much higher than those for drying gases. In practice, EBCT for liquids depends on the contaminant and is usually 2e20 min. A minimum of 6 min per column is typically used for the treatment of contaminated groundwater using granular activated carbon. For drying of some liquids, an empty bed contact time of 1 min or less may be sufficient. EBCT for air dryers vary from 5 to 30 s.

(d) Breakthrough time (tb) The breakthrough time (tb) is the time required to attain the breakthrough concentration cb, the maximum limit that can be discarded/tolerated by the downstream process. This often lies between 1% and 5% of ci, where ci is concentration of adsorbate in the inlet fluid phase. tb is read from the breakthrough curve generated from pilot test data and usually plotted as c/ci against t, the elapsed time, starting from c/ci ¼ 0. Also,   Vb (12.22) tb ¼ Q pilot Where Vb is the cumulative fluid volume at breakthrough and Qpilot is the fluid inflow rate in the pilot plant. The breakthrough time may be considered same for the pilot plant and the scaled up plant when LR (or Us) and adsorbent bed length are the same.

(e) Fraction of bed utilised (f) Considering similar bed characteristics in pilot scale and actual plant, fraction of bed utilized (f) for adsorption is the same for the two cases. This can be estimated as f ð%Þ ¼

ðAdsÞb  100 ðAdsÞex

(12.23a)

Where (Ads)ex is the mass of adsorbate removed at exhaustion and (Ads)b is the mass of adsorbate removed at breakthrough. Typical design value of f as stated above is 85%e90%. f can also be obtained from the breakthrough curve as the ratio of the area (ABDE) between the curve and the concentration axis between t ¼ 0 to tb to the area (ABCDE) between the curve and the concentration axis from t ¼ 0 to t ¼ tex. This is illustrated in Problem 1 (Figure P12.1). Under the assumption that the breakthrough curve is steep, tex is far from tb, and ci >> cb. Also, (Ads)ex ¼ Vexci and (Ads)b ¼ Vbci where Vb and Vex are the respective cumulative fluid volume passed till breakthrough point and exhaustion point (corresponding to cex which is approximately equal to ci) and ci >> cb. This gives, f ð%Þ ¼

Vb  100 Vex

(12.23b)

(f) Adsorbate loading (qs) Mass adsorbate removed per unit weight of adsorbent utilised can be approximated as qs ¼

Q  tex  ci Mads

(12.24a)

374

Chapter 12 Adsorption

It may be approximated from pilot plant data using breakthrough information as: qs ¼

Vex  ci Vpilot  f  rB

(12.24b)

Where Vpilot, is the volume of adsorbent in pilot column and rB is the adsorbent bulk density. qs obtained from Eq. (12.24) may be used to cross check tb obtained from pilot plant data. Using the following expression involving adsorbent usage rate AdUR and the total amount of adsorbent consumed tb ¼

Mbed  f AdUR

(12.25)

Where AdUR ¼ ðQ  ci Þ=qs

(12.26)

(Qci) is the adsorbate loading, ci is inlet adsorbate concentration in actual column and qs is adsorbate loading as obtained from breakthrough curve. Problem 2 illustrates a more accurate estimation of the aforementioned parameters.

Bed design Contactor Dimensions: For gas adsorption, typical L is between 0.3 and 1.2 m, depending on the solute concentration in the gas stream. At times L is selected based on the amount of adsorbent to be provided by the manufacturer, and EBCT is recalculated to check if it is within permissible limits. The final selection of contactor dimensions is made from standard vessel dimensions discussed in Chapter 17. Accordingly, the loading rate (LR) or superficial velocity (Us) is modified and checked to lie within permissible limits. Cycle time: The cycle time comprises of time required for adsorption (tads) which is equal to the breakthrough time (tb) obtained from pilot test, desorption time (tdes), and time required for changeovers. Efforts are made to ensure minimum desorption and change over time as this builds in some cushion in plant capacity and becomes helpful as the adsorbent deteriorates with time. Initially, one can assume tads ¼ tdes, which gives tcycle ¼ 2tads, but this needs to be checked with the actual time for desorption required, in case (EBCT) is different for pilot and actual plant due to adoption of standard contactor dimensions ðtcycle Þ ¼

ðEBCTÞ  ðtcycle Þpilot ðEBCTÞpilot

(12.27)

In practice, cycle time for the dryer is specified between 8 and 24 hr. Mass of adsorbent required (Mads): This may be estimated as follows Mads ¼ VrB It may also be obtained from adsorbate loading using Eq. (12.24) Mads ¼

Q  ci  tads qe

(12.28)

12.2 Packed bed adsorption

375

Considering that only a fraction f of the bed is utilised for adsorption, the adsorbent required is Mbed ¼ Mads =f In order to provide extra nonadsorbent volume as margin for design and operational uncertainty, Mbed is further multiplied by a safety factor SF to give Mbed ¼

Mads  ðSFÞ f

(12.29)

Typically SF is taken as 1.2e2.5. It depends on the designer’s experience with similar systems as well as the reliability of the data used in the specific design.

Volume of fluid treated/change out period Vf ¼ Vb ¼ Qtb

(12.30)

Pressure drop Pressure drop is of extreme importance in fixed-bed adsorber design. The processing rate can be limited by pressure drop as excessive pressure drop results in bed compaction or lifting and very lowpressure drop results in uneven distribution and channeling. Nevertheless, most adsorption plants operate with a small pressure drop across the adsorbent bed to keep power costs low because large particles are used whenever possible and velocity is also typically low to approach equilibrium between fluid and adsorbent. Typical pressure drop through a vapour-phase carbon bed is 8e35 cm of water column per m of bed. Table 12.8 provides the pressure drop guidelines for spherical, granular or extruded adsorbent with 1.5  dp  6 where dp is the nominal particle size (mm).

Table 12.8 Pressure Drop Guidelines for gaseous and liquid services for spherical, granular, or extruded adsorbent with 1.5 £ dp (mm) £ 6. Pressure drop range (cm WC/m bed depth) Bed characteristics

Gas flowing

Liquid flowing

Uneven distribution and channelling

2250

376

Chapter 12 Adsorption

Bed pressure drop is a function of process fluid properties, adsorbent characteristics, and vessel dimension. It is usually estimated by Ergun Equation for both gases and liquids. Dp a3 dp;eff rf 150ð1  aÞ ¼ þ 1:75 L ð1  aÞG2f Rep

(12.31)

Where ðDp=LÞ is the pressure gradient in Pa/m of adsorbent bed, Gf is superficial mass flux in kg/ (s$m2), rf is process fluid density (kg/m3) and a is bed void fraction. Rep, the particle Reynolds number d Gf is defined as Rep ¼ p;eff m , dp,eff is the effective diameter of the particles (m), i.e., the diameter of a f

sphere of same surface to volume ratio as the adsorbent particles which is not usually equal to the nominal particle size and is given by dp;eff ¼ 6 ð1aÞ ap j , where ap$is the specific surface area (particle

external area/volume, m2/m3) of the particle and j is the particle shape factor (1.0 for beads, 0.91 for pellets and 0.86 for flakes). It is important to note that (1) The constants in Ergun Equation are obtained for specific packings and the equation may not be strictly valid for adsorption columns packed with granular or pellet form of particles. (2) The correlation has been obtained for steady-state flow conditions and one needs to be careful while applying them to situations when velocity rapidly changes with time, e.g., in pressurisation and depressurisation steps of PSA process (3) The highest pressure drop is likely to occur during the regeneration step since the fluid is at its highest temperature and/or lowest pressure in this step. It is therefore important to perform pressure drop analysis for each step in the cyclic operation of a fixed-bed adsorber. (4) The pressure drop is often high in low-pressure gas-phase applications like vent gas cleaning and solvent recovery due to low density and high velocity of the fluid. Under this condition, shallow adsorption beds (low L/D ratio) may lead to lower pressure drop. However, adsorption beds with low aspect ratio often result in flow mal-distribution and may require flow distribution systems like manifolds, baffles, and screens, which may add to the pressure drop and offset the advantage gained. Due to this, granular rather than powder adsorbents are used. Use of specially designed adsorbent, such as trilobe and monolith, also assists in keeping pressure drop low in low-pressure gas-phase applications. (5) For situations where superficial velocity rapidly changes with length as in bulk separation processes (air separation), a summation of pressure gradient along the incremental length is necessary to obtain an accurate estimate of pressure drop.

Bed configuration and mode of operation Total feed flow rate, allowable pressure drop, length of MTZ, method of adsorbent regeneration, and capital investment are the factors that determine the number and arrangement of fixed beds. In order to achieve a steady flow of product, the simplest arrangement requires two columns operated simultaneously e one in adsorption and the other in regeneration mode. Regeneration is done when the MTZ approaches the bed outlet, i.e., most of the bed is saturated. However, this is possible only when the breakthrough curve is steep, and the first column can be operated safely until the breakthrough point to discharge effluent substantially free of solute. In this case, the influent stream is diverted to the second column, and the first is set for regeneration.

12.2 Packed bed adsorption

377

When the breakthrough curve is relatively flat so that a substantial amount of adsorbent remains unsaturated at the breakpoint, the gas flows through the second adsorber in series with the first until the adsorbent in the first column is almost entirely saturated, but a breakthrough is not allowed in the lag adsorber. The first bed is then sent for regeneration, and the influent stream is passed through the second bed which now becomes the lead vessel and a freshly regenerated third adsorber in series is now the lag bed. Thus, in practice, more than two beds are often used, which introduces the need for complex piping and valve arrangements together with a control system. While multiple beds in series are preferred for a long MTZ, a series-parallel combination of multiple beds is a likely choice for high flow rates and large MTZ lengths. Multiple beds in parallel is used for a relatively high flow rate and a short MTZ length to substitute a single large vessel difficult to ship and operate. Usually, the beds are identical, and the feed is split equally among them, and the product streams are subsequently rejoined. It is desirable to have a separate regeneration system for each bed to enable independent flow. Usually, flow through a packed bed is in the vertical direction, and a proper flow distribution is ensured by providing adequate plenum space above and below the packed bed. The same effect can be achieved by employing perforated baffle plates, and placing the inlet and outlet nozzles symmetrically. Details on the flow distributor can be found in Chapter 14. Fixed-bed adsorbers are commonly vertical cylindrical vessels. Such an orientation ensures symmetric fluid distribution and is opted for Q < 1.2 m3/s. A typical bed for water treatment may use filter blocks at the bottom covered by smaller sized particles (commonly ceramic and gravels) and have sand as the uppermost layer. When large volumes need to be treated by small amounts of adsorbent, the pressure drop becomes excessive except for shallow beds. Under such conditions, a horizontal bed with vertical flow is preferred. A shallow bed in a horizontal vessel is also selected for low-pressure gas systems like vent stream clean up and recovery of solvent vapour where low density and large velocity of the gas stream causes a high-pressure drop. However, in the horizontal configuration, since settling/bed movement is inevitable, there is always a tendency for fluid to bypass the adsorbent. At times, the challenges involved in uniform flow distribution often offset the savings in pressure drop. In vapour-phase adsorption, the flow direction is usually upward during adsorption to accommodate some bed expansion. During desorption, (usually the limiting step), the flow is in the downward direction as allowable velocities are greater for crushing than lifting. In temperature swing, keeping the desorption flow countercurrent to adsorption leads to the lowest residual loading. The critical parameter for upflow at high flow rates is bed lifting. This occurs at the onset of fluidisation where the pressure drop is expressed as Dp ¼ ð1  aÞðrs  rf Þg (12.32) L The critical consideration for downflow is bed crushing which occurs when the summation of pressure drop and bed weight exceeds the compressive pressure for crushing. While gas adsorption operates in upflow mode, downflow operation is generally preferred for liquid streams. If all the adsorption stages involve liquid, the design considerations are similar to gas phase adsorption. However, when the regeneration step of a liquid-phase adsorbent involves a vapour, a fill

378

Chapter 12 Adsorption

(upflow) and a drain step (downflow) is necessary. It is preferable to refill the adsorbent bed in upflow for an easier displacement of gases and vapours and prevention of maldistribution in the subsequent adsorption step. Refilling should be done for sufficient length of time to ensure complete removal of the gas pockets; otherwise, products in the adsorption step may be contaminated, causing excessive bed lift. The design of the drain is critical to ensure uniform exit of the mass transfer zone from the column. A minimum of 30 min should be provided for thorough draining, and even after careful drainage, the liquid holdup can be 40 cc/100 g adsorbent. The drain may also be used to introduce backwash liquid, and in that case, it should ensure uniform distribution of liquid across the entire bed cross-section. Backwashing is the process of reversing flow through a bed to dislodge material trapped in pores or attached to media. It is essential for liquid streams containing suspended solids. In this case, downflow with backwashing capabilities is the best option, and the column is provided with a 20%e50% bed expansion allowance for adsorbent backwashing before the contactor is put in service. It not only removes solid accumulation but also reduces microbial growth and pressure drop due to fouling of the bed. Sand filters in water service are typical with backwashing facility. Upflow adsorption is preferred for liquids containing suspended solids as the bed, in this case, does not act as a filter and gets plugged by deposited solid. Downflow mode is also not preferred if the equipment is susceptible to biological fouling. The design of an adsorption system for the liquid phase should thus include the contactor as well as accessories, such as distributor, bed support, drain, and backwash equipment. Generally, adsorption is performed at the temperature and pressure of the inlet fluid. In case options are available, the lowest temperature and the highest pressure (for gases) is selected to maximise the adsorptive load. If a gaseous feed contains a condensable vapour component, the temperature (pressure) needs to be high (low) enough to prevent condensation inside the bed. However, for PSA, the adsorption pressure is decided from economics e larger swings lead to better adsorption but consume more compressive power.

12.3 Design illustration These illustrate only the steps of bed design that includes adsorbent quantity, breakthrough time, bed orientation, bed dimensions, and operating conditions. The reader should follow Chapters 14 and 17 to arrive at the details of the adsorber vessel. An adsorber is to be designed for air decontamination. This air stream (w200 m3/hr, 1 atm, contains acetone (MW ¼ 58) vapour (0.13 mol/m3) from a pharmaceutical process. Maximum limit of acetone concentration in the treated air is 0.004 mol/m3. Solution Step 1: Literature survey and selection of adsorbent. Brosillon, Manero, and Foussard (Enviro. Sci. Technol., 2001, 35, 3571e3575) reported column adsorption experiments for acetone removal from air using Zeolite adsorbent. The same adsorbent is considered for design and the experimental data is treated as pilot plant data. Reported experimental details for the plant are e

Problem 1.

20 C)

12.3 Design illustration

379

Adsorbent: Silicalite, a commercial zeolite having MFI structure as trilobe particles. 1 mm

External area

1180 m2/m3

Lobe height

5 mm

Pellet porosity

0.5

Particle density

1143 kg/m3

Tortuosity

4

Si to Al ratio

47 to 70

Lobe diameter

10

10  10

Pore diameter

m

3

Particle diameter for pressure drop estimation, dp,effective ¼ 4  10

m

Adsorption column Bed volume

1L

Bed height

0.2 m

Temperature

298 K

Superficial velocity

0.29 m/s

Bed porosity

0.4

Bed density

4

Inlet air humidity

0

Si to Al ratio

700 kg/m3

Feed: 0.13 mol/m3 acetone mixed in air at 1 atm, 20 C. mfeed ¼ 1.825  105 kg/(m.s) is taken for air as acetone concentration is small. rfeed ¼ 1.20175 kg/m3, calculated from feed composition, temperature and pressure. MWacetone ¼ 58, MWair ¼ 28.8 Humidity of air nearly “zero.” Breakthrough curve obtained experimentally is as shown below e 1.0 E

D

0.97

C

0.8

Area ABCEA = 40.18

0.6 C/Ci

Area ABDEA = 26.98 f = 26.98/40.18 = 0.67

0.4

0.2 0.031 A

0 0

B 20

81.5

28.3 40

60

80

100

120

time (min)

FIGURE P12.1 Concentration breakthrough curve and estimation of bed fraction (f) utilized.

380

Chapter 12 Adsorption

Step 2: Experimental conditions, breakthrough, and bed exhaustion. Superficial velocity in bed, Us,pilot ¼ 0.29 m/s In this problem, with notations used in the chapter text, ci ¼ 0.013, cb ¼ 0.004, cb/ci ¼ 0.0308. We assume bed exhaustion at cex/ci ¼ 0.97, i.e., cex ¼ 0.1261. From the breakthrough curve, the breakthrough time and the bed exhaustion time are read as tb ¼ 28.3 and tex ¼ 81.5 min. In fact, the numerical data available in the paper has been used for more accurate interpolation. Summary of the data pertaining to the pilot plant experiments: Lpilot ¼ 0.2 m, Vpilot ¼ 1000 cm3 ¼ 1 L ¼ 1  103 m3, Qpilot ¼ 2 m3/hr ¼ 5.556  104 m3/s EBCTpilot ¼ Vpilot/Qpilot ¼ 1.8 s Step 3: Design calculations Qplant ¼ 200 m3/hr ¼ 5.556  102 m3/s Fraction of bed utilised up to breakthrough e f¼

ðShaded area in the breakthrough plot from t ¼ 0 to tb ¼ 28:3 minÞ ¼ 26:98=40:18 ¼ 0:67 ðShaded area in the breakthrough plot from t ¼ 0 to tex ¼ 81:5 minÞ

Adsorbate transfer rate in pilot plant per unit total bed volume. ¼ Qpilot  (cicb)/Vpilot ¼ 2  (0.13  0.004)/1 ¼ 0.252 mol/(hr. liter bed) Since 67% of the bed gets utilised, adsorbate transfer rate in pilot plant per unit bed volume utilized, ¼ Qpilot  (cicb)/f ¼ 0.252/0.67 ¼ 0.376 mol/(hr. litre bed) Adsorbate transfer rate in plant. ¼ Qplant  (cicb) ¼ 200  (0.130.004) ¼ 25.2 mol/hr Adsorbent loading per litre of bed volume utilized. ¼ 0.376  tb ¼ 0.376  28.3/60 ¼ 0.1773 mol per litre bed. Considering tcycle ¼ 8 hr, tads ¼ 4 hr, Adsorbate transfer in plant during tads ¼ 4  25.2 ¼ 100.8 mol. Therefore, bed volume that gets utilised in the plant ¼ 100.8/0.1773 ¼ 568.5 L. A safety factor SF ¼ 1.3 is considered as the pilot plant data used is for fresh adsorbent and the scaled up plant will use regenerated adsorbent in cycles. This takes care of inefficient regeneration as well as adsorbent ageing. Adsorbent volume in plant for utilisation up to breakthrough ¼ 1.3568.5 ¼ 738.8 L. Total adsorbent volume to be provided in plant ¼ 1.3  568.5/f ¼ 738.8/0.67 ¼ 1103 L. Mass of adsorbent to be provided in the plant ¼ 1103  700/1000 ¼ 772 kg. Finding Dplant and Lplant L/D to lie between 3 and 4, and hence, assumed Lplant/Dplant ¼ 3. rðp = 4Þ  ðDplant Þ2  ð3  Dplant Þ ¼ 1:103 m3

12.3 Design illustration

381

Dplant ¼ 0.78 m and Lplant ¼ 2.34 m. The diameter lies in the typical range of 0.3e1.2 m for gas service. 1 Superficial velocity in plant Us;plant ¼ ð200 =3600Þ  ¼ 0:12 m/s. ðp=4Þ  ð0:78Þ2 The above is less than 0. 29 m/s (Us,pilot), and hence, OK. Checking for (L /dp,effective) > 100: this is obviously met. The other limit of Dplant / dp,effective is met too. Checking for adequate EBCT: This is required to ensure sufficient time of contact for the desired rate of solute transfer. EBCTplant ¼ Vplant =Qplant ¼

1103  103 ¼ 19:9 s ð200=3600Þ

EBCTplant is more than EBCTpilot (¼1.8 s). It is also comparable with 30 s, which is a typical limit for a well-designed air drier. Other performance parameters Feed processed per cycle ¼ Qplant  tads ¼ 2004 ¼ 800 m3 at 1 atm, 20 C Adsorbate in bed before regeneration ¼ 800  (0.13  0.004)58 ¼ 5865 g Rep ¼ (dp,effective  Us,plant  rfeed)/mfeed ¼ (4  103  0.12  1.20175)/1.825  105 ¼ 31.61, this satisfies the 20e40 limit. Orientation: Vertical, since the feed flow rate is below 1.2 m3/s. The adsorbent bed shall be sandwiched between two beds of 6 mm f ceramic balls, each 150 mm deep, to ensure uniform flow distribution and bed stability. This would make the total bed length 2  0.15 þ 1.34 ¼ 1.64 m. Bed pressure drop is given by Eq. (12.31). Rep ¼ 31.61; a ¼ 0.4; dp,eff ¼ 4103m; rf ¼ 1.20175 kg/m3, and Gf ¼ rf  Us,plant ¼ 1.20175  0.12 ¼ 0.14421 kg/(m2$s). L ¼ 1.34 þ 0.3 ¼ 1.64 m, the pressure drop (Dp/L) in the larger size ceramic balls of 0.3 m total depth will be lower than that in the adsorbent bed. Hence, using L ¼ 1.64 m for pressure is a conservative (lower) estimate. Dp/L ¼ 496.1 Pa/m and Dp ¼ 1.64  496.1 ¼ 813.6 Pa. Bed pressure drop Dp (813.6 Pa) is small compared to the total pressure (1 atm ¼ 101.3  103 Pa), and hence, the consideration of the design pressure of 1 atm is OK. The pressure drop at inlet nozzle, outlet nozzle, and distributor (if provided) is also added to the bed drop. The said components are to be sized/designed so that the total pressure drop, including all components, remains low compared to 1 atm. Design output summary Adsorbent bed Adsorbent: Zeolite adsorbent, trilobe particle shape, 4 mm effective particle diameter. Amount of adsorbent in bed: 1103 L (772 kg) Length: 1640 mm of adsorbent packing including 2  150 mm of 6 mm diameter ceramic balls upstream and downstream.

382

Chapter 12 Adsorption

Diameter: 780 mm f. Bed orientation: vertical. Operation Feed rate: 200 m3/hr, at 1 atm, 20 C Cycle time: 8 hr, Adsorption time per cycle: 4 hr. EBCT for the adsorbent bed 19.9 s. Cumulative volume of feed processed per cycle 800 m3, at 1 atm, 20 C. Bed pressure drop: 813.6 Pa, including ceramic ball packed depth but excluding pressure drop in inlet and outlet nozzles, and distributor, if provided. A 200 m3/day stream of wastewater containing 200 mg/L of total organic carbon (TOC) needs to be treated to less than 10 mg/L TOC content. Design the adsorption column based on a column (9.5 cm diameter. 175 cm long, 60 L /hr inlet flow rate) study with a cheap 8þ20 mesh (0.841e2.38 mm) coal-based adsorbent (Ref. Table 12.6) having bulk density 500 kg/m3. Cumulative volume processed up to breakthrough and bed exhaustion are: Vb,pilot ¼ 8500 L; Vex,pilot ¼ 9600 L. Solution Data

Problem 2.

ci ¼ 200 mg/L; cb ¼ 10 mg/L; Dpilot ¼ 9.5 cm, Lpilot ¼ 175 cm. 2   Vpilot ¼ ðp =4Þ  9:5  102  175 102 ¼ 0:012404 m3 ¼ 12.404 L. Vb,pilot ¼ 8500 L; Vex,pilot ¼ 9600 L; Qpilot ¼ 60 L/hr ¼ 1 L/min; tb,pilot ¼ Vb/Qpilot ¼ 8500/1 ¼ 8500 min; tex,pilot ¼ Vex/Qpilot ¼ 9600/1 ¼ 9600 min Qplant ¼ 200 m3/day ¼ 200  103/(24  60) ¼ 138.9 lpm ¼ 0.1389 m3/min Particles: 20/8 mesh, i.e., 0.841e2.38 mm, j ¼ 1, assumed as the particles are small. Average particle size, dp ¼ (0.841 þ 2.38)/2 ¼ 1.60 mm; Since the feed is dilute, we take density and viscosity same as water at 27 C, i.e., rf ¼ 0.9965 g/cc ¼ 996.5 kg/m3; mf ¼ 0.8591103 Pa s. Qpilot 1 ¼ ¼ 141:08 lpm/m2, this is within the typical LRpilot ¼ 2 2 ðp=4Þ  ðDpilot Þ ðp=4Þ  ð9:5  102 Þ limits of 80 and 240. Keeping the same LR, scaled up plant cross-section, Aplant ¼ Qplant =LR ¼ 138:89=141:08 ¼ 0:9845 m2 rDplant ¼ ðAplant  4=pÞ1=2 ¼ ð0:9845  4=pÞ1=2 ¼ 1:12 m EBCTpilot ¼ Vpilot =Qpilot ¼ 12:404=1 ¼ 12:404 min EBCTplant is kept the same as pilot scale to ensure that the scaled up plant adsorbate transfer rate is close to that in the pilot scale. This is between 6 and 20 min, the range for liquid-phase adsorption in an industrial scale. Based on EBCTplant ¼ 12.404 min. Lplant ¼

EBCTplant  Qplant ðp=4Þ  ðDplant Þ

2

¼

12:404  138:9  103 ðp=4Þ  ð1:12Þ2

¼ 1:75 m

12.3 Design illustration

383

Vplant ¼ EBCTplant  Qplant ¼ 12:404  138:9  103 ¼ 1:723 m3 Assuming symmetric breakthrough curve in its rising zone in pilot plant, the adsorbate transfer rate may be estimated as: Adsorbate transfer from tb to tex ¼ (Vex  Vb)  (cb þ ci)/2. A conservative estimate of adsorbate transfer rate in pilot plant is based on effluent concentration always close to cb. rAdsorbate transfer (from t ¼ 0 to tb) ¼ Vb(ci  cb) f ¼ ¼

Adsorbate transfer from t ¼ 0 to tb Adsorbate transfer from t ¼ 0 to tex

Vb  ðci  cb Þ 8500  ð200  10Þ ¼ 0:93 ¼ Vb  ðci  cb Þ þ ðVex  Vb Þ  ðci þ cb Þ=2 8500  ð200  10Þ þ ð9600  8500Þ  ð200 þ 10Þ=2

In a real plant, there are fluid distribution problems and f is usually between 0.85 and 0.9. Hence, we assume, f ¼ 0.87 for the scaled up plant as a conservative estimate for design. Adsorbent loading in the utilised portion of pilot plant bed at breakthrough, based on f ¼ 0.93, Vb;pilot  ðci  cb Þ 8500  ð200  10Þ ¼ 139887 mg=L ¼ 12:414  0:93 Vpilot  f Same loading is assumed in the plant, i.e., qs,plant ¼ 139887 mg/L of the utilised portion of bed at breakthrough. Such a bed is expected to be run continuously for 5 days or more, and hence, assumed e tads,plant ¼ 4 days. Adsorbate transfer during this period qs;pilot ¼

¼ tads;plant  Qplant  ðci  cb Þ ¼ 4  200  103  ð200  10Þ ¼ 152  106 mg r Utilized bed volume in the plant bed at breakthrough ¼ (152  106)/139887 ¼ 1086.6 L. Considering f ¼ 0.87, total bed volume ¼ 1086.6/0.87 ¼ 1249 L. Considering a safety factor, SF ¼ 1.5, total bed volume, ¼ 1.5  1249 ¼ 1873.5 L. This bed volume is little above the bed volume estimated earlier considering EBCT, and is, therefore, OK. Reestimation of Dplant and Lplant based on L/D ¼ 3 for the adsorbent packed bed Vplant;revised ¼ ðp = 4Þ  ðDplant Þ2  ð3  Dplant Þ ¼ 1:8735 m3 ; Dplant, revised ¼ 0.93 m, Lplant, revised ¼ 30.93 ¼ 2.79 m, say 2.8 m. rFinal dimensions Dplant, revised ¼ 0.93 m, Lplant, revised ¼ 2.8 m, Vplant;revised ¼ ðp =4Þ  ðDplant Þ2  Lplant ¼ 1:9 m3 Qplant 138:9 LR ¼ ¼ ¼ 204:5 lpm/m2, well within limit of 240. 2 ðp=4Þ  ðDplant Þ ðp=4Þ  ð0:93Þ2 Mass of adsorbent in bed, Mads ¼ Vplant, revised  rbulk ¼ 1.9  500 ¼ 950 kg. 0:87  1:9  103  139887 ¼ 6:1, say 6 days. This is higher than Breakthrough time of plant ¼ 200  103  ð200  10Þ the original assumed value but is reasonable.

384

Chapter 12 Adsorption

Us;plant ¼ Rep ¼

Qplant 200=ð24  60  60Þ ¼ 3:41  103 m=s 2 ¼  ðp=4Þ  ð0:93Þ2 ðp=4Þ  Dplant;revised

dp;effective  Us;plant  rfeed 1:6  103  3:41  103  996:5 ¼ ¼ 6:32 0:8591  103 mfeed

Bed orientation, support, pressure drop, etc. Orientation: Vertical. The 930 mm f adsorbent bed shall be sandwiched between two beds of 4 mm f gravel, each 150 mm deep, to ensure even flow distribution and bed stability. This would make the total packed bed length 2.8 þ 0.3 ¼ 3.1 m. Bed pressure drop is given by  2 Dp ð1  aÞGf 150ð1  aÞ ¼ 3  þ 1:75 L Rep a dp;eff rf Rep ¼ 6.32; a ¼ 0.4 e although this data is not provided, the voidage is expected to be close to 0.4 as the ratio of bed diameter to particle diameter increase beyond 15, which is true in this case. dp,eff ¼ 1.6103m; rf ¼ 996.5 kg/m3, and   Gf ¼ rf  Us;plant ¼ 996:5  3:41  103 ¼ 3:4 kg= m2 $ s ; L ¼ 3.1 m, the pressure drop (Dp/L) in the larger size gravel of 0.3 m total depth will be lower than that in the adsorbent bed. Hence, using L ¼ 3.1m for pressure is a conservative (lower) estimate. rDp L ¼ 1087 Pa/m and Dp ¼ 3.11087 ¼ 3370 Pa. Bed pressure drop Dp (3370 Pa) is small enough. The inlet nozzle, outlet nozzle, and the distributor (if provided) pressure drop is also added to the bed pressure drop to obtain the total pressure drop across the adsorber vessel. The said components are to be sized/designed so that the total pressure drop is within the allowable limit. Typical maximum limit for similar applications is w0.5e0.7 atm, and there is sufficient margin for adequate sizing of the nozzles. Design output summary Adsorbent: Coal-based adsorbent particles, 8 þ20 mesh. Amount of adsorbent in bed: 1900 L (950 kg) Length: 2800 mm of adsorbent packing between 2  150 mm of 4 mm diameter gravel. Diameter: 930 mm. Bed orientation: vertical, downflow during adsorption. Pressure drop: 3370 Pa.

Further reading Crittenden, B., & John Thomas, W. (1998). Adsorption technology and design (1st ed.). Elsevier. Treybal, R. E. (1980). Mass-transfer operations (3rd ed.). New York: McGraw-Hill.

CHAPTER

Extraction

13

13.1 Introduction Liquideliquid extraction (LLE), also termed as solvent extraction, is a separation process where component(s) from a solution is recovered by the addition of a second immiscible liquid (solvent) which has a greater affinity for the component(s) to be recovered. As mentioned in Chapter 9, the solution left after extraction is termed raffinate and the solvent-rich phase containing the solute is termed extract. Similar to other mass transfer operations, the process is carried out by bringing the feed solution and solvent into intimate contact for a sufficient length of time to ensure maximum solute transfer and subsequently, allowing the phases to separate, generally by gravity. Compared to absorption and distillation, liquid-liquid phase separation is more difficult and slow since the density difference between the phases is not large. Centrifugal force is used for phase separation in case of closer density of the liquid phases. The solute (desired product) is then recovered from the solvent usually by distillation and the solvent is recycled along with fresh makeup solvent to compensate for the losses. Extraction processes therefore always require additional steps to recover and recycle the solvent and the operating and capital cost of solvent recovery is usually higher as compared to the extraction step. Therefore, extraction is opted as a separation process usually when other mass transfer operations namely distillation is (a) infeasible or (b) expensive or (c) requires complex sequencing and when opted, it is important to consider the solvent recovery aspect at the design stage itself. The conditions under which extraction is preferred include the following: (i) Solution of components with close relative volatility, e.g., separation of acetic acid from aqueous solution using methyl tertiary-butyl ether (MTBE) as solvent which can be evaporated easily to recover the acid. Extraction of aromatic compounds from lube oil vacuum distillation cuts using solvents like phenol, furfural, n-methyl pyrrolidone (NMP) is a standard practice. (ii) Azeotropic mixtures particularly when the azeotrope cannot be split by adding a third component or a change in the operating pressure does not influence the volatility of one component more than the other. Typical examples include separation of tetrahydrofuran, pyridine and formic acid from water or separation of dichloromethane or ethyl acetate from ethanol. (iii) Heat-sensitive products encountered in food, pharmaceutical and green chemistry biomolecules namely vitamins, penicillin, flavours and fragrances as well as certain aldehydes and organic acids in chemical industries. Process Equipment and Plant Design. https://doi.org/10.1016/B978-0-12-814885-3.00013-0 Copyright © 2020 Elsevier Inc. All rights reserved.

385

386

Chapter 13 Extraction

(iv) Nonvolatile metal ions (precious or rare earth metals) in an aqueous solution by addition of an organic solvent, e.g., separation of copper/iron, nickel/cobalt or chromium/vanadium. In most cases, one of the ions forms a chelate complex in the organic phase, while the other remains in the aqueous solution. (v) Nonvolatile organic compounds (higher boiling point than water) from aqueous effluents, e.g., removal of phenols, cresols, aniline or other aromatic derivatives from industrial waste water, removal of organic compounds during oxidation of organic products, production of caprolactam, etc. In such cases, a low boiling point solvent with a higher concentration of solute makes extraction economically more feasible than distillation. For separation of high molecular weight fatty acids from vegetable oils, vacuum distillation is more expensive and the preferred option is extraction using liquid propane as solvent.

13.2 Extractor types and selection The classification of extractors based on different criteria is presented in Fig. 13.1. These are classified based on (a) the mode of contacting of the two liquid phases and (b) the nature of solvent. Extractors may use a single solvent, thus forming ternary systems or systems equivalent to ternary systems where two components are to be separated by addition of the third component. A mixed solvent system employs a solvent solution of at least two components where the solubility relationship cannot be reduced to an equivalent ternary system. Double solvents are used for fractional extraction where the solution to be separated is distributed between two immiscible solvents, thus comprising a fourcomponent system. In this chapter, we limit our treatment only to extraction with single solvents.

13.2.1 Extractor types Similar to other fluidefluid operations, the two phases can be either in continuous contact, separated only at the end of the operation, or be contacted in stages, i.e., contacted and separated sequentially where the number of contacts constitutes the number of stages (Fig. 13.2). In an ideal stage, the two liquids are assumed to attain equilibrium during each contact. In continuous contact, equilibrium is not normally achieved. Both types of equipment with continuous or stagewise contacting can be operated continuously.

Stagewise contact Flow arrangements of the different contacting schemes along with the associated solvent recovery sections are shown in Fig. 13.2. Simple case of a single stage of contacting with sections for solvent recovery (SR) from extract and raffinate streams is shown Fig. 13.2A. In multistage cases, the fluids can be in cross or countercurrent flow. In cross flow, raffinate is contacted repeatedly with fresh solvent (Fig. 13.2B). This is generally uneconomical for large commercial processes and is used in laboratories to generate equilibrium data and to demonstrate feasibility of solute removal to low residual concentration. Countercurrent operation (Fig. 13.2C) requires less solvent for same feed rate, solvent

Extractor

Continuous (Countercurrent flow)

Mode of contact

Discrete (Stagewise)

Mixersettler

Cross flow

Multistage

Spray column

Countercurr ent flow

Countercurr ent flow with reflux

Packed column

Centrifugal extractor

Sieve tray

Wetted wall column

Agitated tower

Karr column

FIGURE 13.1 Classification of extractor types.

Scheibel column

Single solvent

RDC

Double solvent

Pulsed column

Mixed solvent

13.2 Extractor types and selection

Single stage

Static tower

Nature of solvent system

387

388

Chapter 13 Extraction

and raffinate composition, forms less amount of extract solution and achieves higher solute concentration in extract. Thus, similar to other mass transfer operations, countercurrent operation is the preferred option. The improvement is less significant for five or more number of stages. Typical stage efficiency is 75%e100% and in most industrial plants, the reduced efficiency is due to incomplete separation (settling) of phases rather than due to inadequate mixing. Continuous countercurrent operation with multiple contacting can be supplemented with a reflux arrangement to enhance recovery. In this case, the extract and raffinate are in countercurrent flow through the cascade of stages with feed solution entering at an intermediate stage and solvent introduced at one end. Reflux may be provided at both ends of the cascade as shown in Fig. 13.2D or only at one end corresponding to the enriching or stripping end of distillation. In some columns, the two end temperatures are sometimes maintained at different values to alter the equilibrium in favour of a purer stream. Temperature control is achieved by a pump back reflux/circulating reflux appropriately located close to one end of the tower. Extraction towers for aromatics from lubricating oil stocks often employ this. While single-stage contact (Fig. 13.2A) can be treated analogous to flash distillation, countercurrent multiple contact is akin to gas absorption and its operation with reflux is analogous to rectification. The fundamental balance equations of the analogous mass transfer operation are similar to the different stagewise operations and the design equations formulated therein can be used, rewritten in terms of flow rates and composition of raffinate (R) and extract (E). However, the equilibrium relationships are more complex due to the mutual solubility of the two liquid phases. As a result, simplifying assumptions like constant molar flow rate and linear equilibrium and operating curves are not directly applicable. Stagewise contacts generally take place in mixer settlers, the details and design of which is discussed in Section 13.5.

Continuous contact Continuous contacting devices have gained importance with the increased application of extraction in petrochemical and chemical industries processing high volumetric flow rates. These towers contain an equivalent of several stages where the two fluids flow countercurrent to each other. In centrifugal extractors, centrifugal force is used for phase separation and is opted for closer density of the extract and raffinate phase. The lighter liquid is introduced from the tower bottom and the heavier phase from the top. Either of the phases can be the dispersed phase, depending upon the proportion of the two phases, interfacial tension and relative viscosity. Usually one liquid is pumped through the equipment and that fixes the maximum velocity of the second liquid (flooding condition). In all commercially important towers, the interfacial area is in the form of droplets where one or both liquids are dispersed by flow through nozzles, orifices, screens, packings, etc., or by agitation. One liquid flowing as a film over a surface while in contact with the second liquid may occur over part of the equipment when the tower internals are preferentially wetted by the dispersed phase. Continuous contact is effected in vertical static or agitated towers. Static towers are opted for low to medium interfacial tension (10e15 dyne/cm) when only a few theoretical stages are required. Packed

13.2 Extractor types and selection

(A)

R F S

SR

389

R' E'

E

SR SE

SR

Recycled S Smakeup

∑E

(B) E2

E1 F

R1

1

2

S1

SE R'

EN

RN-1

R2

E'

SR

N

RN

SR

SN ∑S

S2

SR

Smakeup

(C)

F(R0) R1

E'

SR

1

R2

RN-1

E3

EN

E2

SR

N

2

E1

RN S(EN+1)

R'

SR Smakeup

SE

(D) SE E1 SR E' R0 PE'

E2 Ee 1

R1 R

Ee+f e

e-1

Re

Ef

Ef+t f

Rf-1 F

Rf

Et Rt-1

Et+f t

EN

EN+1 N

Rt

RN-1

SM

R N RN PR SR

Smakeup

SR

N' PR'

FIGURE 13.2 Stagewise contacting configurations with solvent recovery. (A) Single stage, (B) crosscurrent multiple stage, (C) countercurrent multiple stage and (D) countercurrent multiple stage with extract and raffinate reflux. Nomenclature used in Fig. 13.2: SR, solvent recovery unit; SM, solvent mixer unit; F, feed solution; S, solvent; R, raffinate; E, extract; subscripts on individual streams denote stage numbers from which the stream exits; R 0 , solvent-free raffinate after solvent recovery unit; E 0 , solvent-free extract; SR , recovered solvent from raffinate; SE , recovered solvent from extract; Smakeup , makeup solvent; PR , saturated raffinate product; PE , saturated extract product; f,e,t, feed stage, typical enriching and a typical stripping stage. SR denotes section of plant for solvent recovery.

390

Chapter 13 Extraction

towers with random or structured packings and sieve plate towers with or without downcomers are commonly used. Spray columns are not very common in industries and are used for rapid irreversible chemical reactions like neutralisation of waste acids. Wetted wall columns are rarely used. Their only advantage is that they do not have any internals. Bubble caps generally display low capacity and low tray efficiency and so are not used for extraction process. Agitated towers are more economical when more than two to three theoretical stages are required and interfacial tension is moderate to high. The different types of agitated towers are Karr tower, Scheibel tower, pulsed tower and Rotating Disc Contactor. In each of these, some form of motion is imparted to the liquid contents for enhancing capacity and mass transfer rate. The motion also provides an additional parameter for operational optimisation, albeit at an increased capital cost. Pulsing during flow can be achieved in both packed and sieve plate towers by connecting the fluid space to a reciprocating plunger pump, bellows pump or high pressure air pulse. Pulsing in plate towers is shown in Fig. 13.3A. The rapid oscillatory motion of short amplitude imparted to the fluid increases mass transfer efficiency and improves radial distribution but there is a small reduction in throughput. In Karr reciprocating plate tower, the sieve trays are imparted an up and down motion. Scheibel tower operates as a series of mixer-settler units where each unit has a rotating turbine agitator to disperse droplets. The droplets coalesce while passing into the outer settling zone as shown in Fig. 13.3B. Rotary disc contactors (Fig. 13.3C) are divided into compartments by horizontal doughnut-shaped or annular baffles and within each compartment, agitation is provided by the rotating motion of a centrally located horizontal disk. Generally the disk is smooth and flat with diameter less than the opening in the stationary baffle.

13.2.2 Contactor selection A suitable contactor for a given extraction process is selected based on (A) liquid physical properties namely density difference and interfacial tension and (B) difficulty of separation denoted by the number of theoretical stages (NTS) to achieve the desired raffinate and extract composition. A preliminary guideline suggests the following: (i) (ii) (iii) (iv)

Mixer-settler batteries for high NTS (2.5  NTS 9) and easy phase separation. Static columns for easy separation with low NTS. Centrifugal contactor for difficult phase separation and relatively low NTS. Agitated columns (with reciprocating, rotating, pulsing devices) for moderate ease of phase separation and high NTS (1.5e9).

Other considerations guiding contactor selection can be obtained from Table 13.1 that summarises the limitations and advantages of the commonly used extractors.

13.2 Extractor types and selection

(A)

391

Lighter phase

Heavier phase

Compressed air

Timer controlled valves

Air

Liquid Lighter phase

Air Bleed Interface control

Interface

Heavier phase

(B)

(C) Roter Lighter phase

Heavier phase

Mixing region Settling region

Baffle Turbine impellor

Shaft Lighter phase

Heavier phase

Fixed/stator baffles Rotating disc(s) h

Lighter phase

Interface Interface control

Heavier phase

Lighter phase

DS Dr D

Interface

Interface control

Heavier phase

FIGURE 13.3 Agitated extractors with the heavier phase dispersed. (A) Pulsed extractor; (B) Scheibel tower and (C) Rotary disc contactor.

392

Chapter 13 Extraction

Table 13.1 Advantages and limitations of common extractors. Equipment

Advantages

Limitations

General applications a

Mixer settler

Large flow Intense mixing promotes mass transfer Less headroom Rapid restart but does not reach steady state quickly Low capital cost

Long residence time , large floor space, not very easy to scale-up, cannot tolerate solids, higher holdup volume, low throughput, large solvent inventory and losses

Metal industries

Spray extraction tower

Cheap Simple to construct Easy to clean Trouble-free operation High throughput

High HETP (3e6 m)

Rarely used Adopted for rapid irreversible chemical reactions like neutralisation of waste acids

Packed extraction tower

High throughput Easy operation and maintenance (no moving parts) Simple operation even at high temperature and pressure conditions Can handle corrosive liquids by proper choice of packing materials

Not suitable for fouling service. Although more efficient than spray tower, backmixing results in a higher HETP compared to pulsed and mechanically agitated towers

Extensively used in solvent refining of lubricating oils, removal of hydrogen sulphide from petroleum fraction, sweetening of naphtha, removal of phenols from ammoniacal liquor, solvent refining of vegetable oils and chemical recovery in synthetic organic chemical industries (operation generally limited to Dr > 30  50 kg/m3, 0:5  a  5, s < 10 dyne/cm and NTS 10)

Sieve tray tower

High capacity Good efficiency (minimum backmixing)

Affected by changes in wetting characteristics

Refining, petrochemicals

Centrifugal extractor

Short residence time Low headspace Moderate floor space Easy to scale-up Low holdup volume Corrosive fluids Moderate capital cost High throughput Rapid restart and reaches steady state quickly

Cannot tolerate solids High-speed device require maintenance Susceptible to fouling and plugging due to small clearance

Pharmaceutical industry especially for low density difference between phases

13.2 Extractor types and selection

393

Table 13.1 Advantages and limitations of common extractors.dcont’d Equipment

Advantages

Limitations

General applications

Pulsed towers

Reliable High throughput Provides an additional parameter for operational optimisation Can tolerate solids Low floor space but useful for liquids of high interfacial tension (30e40 dyne/cm).

Not easy to scale-up High capital cost Large building headroom Limited stages due to backmixing Limited diameter/height due to pulse energy required

Best suited for nuclear applications due to lack of seal Suited for corrosive applications

Karr reciprocating plate towers

Highest capacity Good efficiency Good turndown capacity (4:1) More uniform droplet size

High capital cost

Chemicals, petrochemicals, refining Pharmaceutical (for difficult systems which emulsify and/or flood early, fouling applications, solids precipitation)

Scheibel tower

High efficiency More uniform droplet size coupled with better phase separation Good turndown capacity (4:1) High flexibility

High capital cost Not recommended for highly fouling/high emulsifying systems

Chemicals, petrochemicals, refining, Pharmaceutical (best suited when many stages required or for low mass transfer rates)

Rotating disc contactor

Suitable for viscous materials (>100 cp) Suitable for fouling materials

High capital cost Limited efficiency (axial backmixing)

Furfural extraction of lubricating oil Desulphurisation of gasoline Phenol recovery of wastewater (suitable for mass transfer controlled systems with few theoretical stages)

Advantageous for processes with relatively slow reactions; a is the volume phase ratio of dispersed to continuous phase; s is the interfacial tension.

a

394

Chapter 13 Extraction

13.3 Choice of solvent Solvent selection is based on equilibrium, economic and environmental considerations. To understand equilibrium considerations, the ternary plot (Fig. 13.4A) that has been discussed in Chapter 9 is considered for extraction of solute B from a feed solution of A and B using pure solvent S. The feed solution has xB;F % B corresponding to point F on the AB arm of the diagram. Addition of adequate amount of pure solvent S shall result in mixture composition corresponding to point M. M shifts towards S with addition of more and more amount of solvent and its locus is the line joining the feed and the solvent composition, i.e., line SF. M inside the two-phase envelope splits into the extract (E) and raffinate (R) phases with composition of B as yB and xB in E and R respectively. yB ; xB correspond to the two ends of the tie line passing through M. Now on considering complete recovery of solvent from the phases E and R, the solvent-free extract and raffinate compositions would be the points E0 and R0 on line AB. A larger distance between points E0 and R0 denotes a greater extent of separation by extraction. One may note that this also depends on the proportion of feed and solvent mixed as the composition of E0 and R0 denotes the composition of the extract and raffinate streams obtainable by using the solvent amount that leads to the feed plus solvent composition at M. The above information is used in selection of solvent. The same construction is done for solvents being compared, keeping the composition same for point M, i.e., the ratio of feed and solvent is kept same. The composition of E0 and R0 is compared and the solvent with a larger distance between E0 and R0 is a better solvent with higher selectivity. Fig. 13.4B shows the equilibrium curves xPz and x0 P0 z0 for two solvents where the former has a larger envelope of immiscibility. The line joining the feed composition xB;F with the pure solvent vertex denotes the locus of the mixture composition that would result if pure solvent is used for extraction. This must lie within the two-phase envelope to create the extract and raffinate phases. Minimum amount of solvent required for the mixture to touch the equilibrium line corresponds to the point Q and Q0 for the two solvents. Relative to Q, the location of Q0 is closer to the solvent vertex S, denoting a higher proportion of S in the mixture for the same feed composition. This illustrates that smaller envelope of miscibility requires a higher amount of solvent for the same separation. Such ternary plots denoting different immiscibility ranges are obtained not only for different solvents but also for the same solvent at two different extraction temperatures since greater miscibility (smaller envelope) results at higher temperatures. However, the advantages of operating at a higher temperature with lower viscosity and mass transfer rate needs to be weighed against higher amount of solvent requirement. Additional factors influencing solvent selection include -

-

Low interfacial tension which is desirable for easy dispersion of the two liquid phases. However, too low a value may lead to stable dispersion and make separation difficult. Since data on interfacial tension are not always available, a rough estimate can be obtained from a difference in surface tension of the two liquids and this is further lowered by presence of emulsifying agents. The presence of minute dust particles also prevents droplet coalescence as they usually accumulate at the interface of an immiscible liquid system. Significant density difference with feed and raffinate throughout the entire range of operation ensures easy separation of extract from raffinate and increases the capacity of the contacting equipment.

13.3 Choice of solvent

(A)

395

B 100

20

80

E' 60

40 F xB,F 60

40

P

E,yB M

80 R'

20

R,xB 100 A

20

40

60

S 100

80

B

(B) 80

20

60

40 P

40

60

P'

20

80

xB,F Q Q' S

A x x'

20

40

60

z'

z

FIGURE 13.4 Ternary plots depicting (A) extraction process and (B) effect of solvent miscibility on extraction.

- Low viscosity requires lower pumping power, leading to higher extraction rate, higher heat transfer and greater ease of handling. Since separation of solvent from feed solution occurs both by gravity and coalescence of dispersed phase droplets, settling is slower for higher viscosity of the continuous phase and smaller drop size. - No azeotrope formation and high relative volatility with respect to the extracted phase within the range of operation is desirable if the solvent is recovered from extract by distillation. Usually in extraction, since the quantity of solvent is larger, choice of a more volatile solute would

396

-

-

Chapter 13 Extraction

require lower heat of vaporisation and lower the operating cost. In case of a highly selective solvent, the amount of solvent may be less and vaporising the solvent is a cheaper option. Low vapour pressure and freezing point enables easy storage and handling of the solvent at atmospheric or moderately high pressure. However, low vapour pressure should not conflict with the requirement of high relative volatility with feed solution that helps separation by distillation. Chemically inert to feed solution and stable under operating conditions. Nontoxic, nonflammable and noncorrosive nature of solvent as well as low cost and easy availability.

In mixer-settler arrangement, liquid with higher volume fraction tends to be the continuous phase. Partial or arrested inversion may result in dual emulsions where the continuous phase is dispersed as droplets in the drops of the dispersed phase. In continuous contact devices, the phase with higher flow rate Selection of dispersed phase is dispersed in sieve tray and packed towers while in all other towers, the liquid with lower flow rate forms the dispersed phase. Presence of interphase at top signifies lighter liquid dispersed and with heavier liquid dispersed, the interphase is at the bottom of the column. Other factors affecting choice of dispersed phase are e

e e e

More viscous phase dispersed for higher capacity (droplet settling/rise slower in viscous liquid) and less viscous liquid dispersed for efficient operation (slow diffusion inside viscous droplets). Emulsifying agents may be used to increase diffusion in viscous droplets. Preferable direction of mass transfer is from continuous to dispersed phase. Phase that preferentially wets the packing material/tower internals is the continuous phase. Dispersion of inflammable liquid is a safer option.

13.4 Design of continuous countercurrent contactors Extraction involves the complex phenomena of droplet breakage and coalescence as well as axial and radial mixing. Presence of impurities significantly affects the interfacial phenomena and equipment performance. As a result, there is lack of reliable data on mass transfer coefficient and interfacial area which makes estimation of process parameters from fundamental theory difficult. The best option is pilot plant testing with the same equipment type selected based on process considerations of the actual production. This is followed by certain empirical scale-up procedures based on experience. The success of design depends on proper solvent selection as well as pilot plant tests with actual liquids for the entire process including solvent recovery over a wide range of operating conditions. Tests with synthetic mixtures resembling plant fluids often lead to unsatisfactory performance. Typical dimensions of pilot columns (based on experience) for different continuous contact extractor types are listed in Table 13.2. The design parameters to be estimated in each case are also mentioned in the table. Scaling-up needs to be based on proven techniques with use of proper safety factors. The generalised scale-up relationships relate column diameter D, volumetric throughput Q, and height per unit stage efficiency HETP as

13.4 Design of continuous countercurrent contactors

Dplant Qplant ¼ k1 Dpilot Qpilot

397

!m1 (13.1a)

and HETPplant Dplant ¼ k2 HETPpilot Dpilot

!m2 (13.1b)

In case of agitated columns, there is an additional parameter f , the frequency of agitation at same specific power input that is related to column diameter. !m3 fplant Dplant ¼ k3 (13.1c) fpilot Dpilot In Eq. 13.1, k1,m1 are capacity scale-up factors, k2,m2 efficiency scale-up factors, and k3,m3 power scale-up factors. Continuous contact equipment is further sized with the active zone height, together with the top and bottom settling zones and instrumentation for control of the column interface. Table 13.2 Typical pilot column dimensions. Column type Dimension

Packed

Sieve tray

Karr

Scheibel

D

50e150 mm

100e150 mm

25 mm

75 mm

Height per theoretical stage

Packing height 2e5 m

1200e1500 mm tray spacing

30e900 mm

3e6 actual stages (75e150 mm)

Process factor

NTS, S/F

NTS, S/F

NTS, S/F

NTS, S/F

Column variable

D, H

D, H

D, H

D, H

Process variable

FþS

FþS

F þ S, f

F þ S, f

Design parameters

F e feed flow rate, S e solvent flow rate, f e frequency of agitation or vibration, H e column height, D e column diameter

Unlike absorption and distillation, processing capacity of extraction equipment (column) refers to the sum of the flow rates of the dispersed and the continuous phase (denoted by subscripts d and c, respectively) and tower cross section is determined from total flooding velocity ðUd þ Uc Þflooding. Prediction of total flooding velocity involves uncertainties and hence extraction columns are designed to operate at 50% flooding velocity or even lower in many cases. The steps of design are as follows: e e e

Choice of solvent, compilation of property data over the operating range Feasibility test and equilibrium data from laboratory tests Contactor selection

398

e e e e e e

Chapter 13 Extraction

Deciding on the dispersed and continuous phase Pilot plant tower design and tests to obtain mass transfer data Calculation of NTS or transfer units and column height Estimation of tower diameter from flooding considerations Design of actual column using scale-up relationships Hydrodynamic calculations to estimate pressure gradient

Flooding Flooding in extractors occurs by three mechanisms e phase inversion due to high flow rate of dispersed phase, entrainment of dispersed phase droplets at high flow rate of continuous phase and presence of contaminants at the interface that causes phase inversion by interface instability. Thus flooding in extractors occur with increase in flow rate of either dispersed or continuous phase and both phases leave from the outlet of the continuous phase. It can be visualised by the appearance of a second interface inside the column. Increase in agitation speed and decrease in interfacial tension promotes flooding.

13.4.1 Calculation of the number of stages In a multistage extraction process, the number of theoretical stages N decides the length of liquid travel and governs the compromise between equipment size or the number of mixer-settler contactors (capital cost) and the ratio of solvent to feed flow rate required for the desired extent of extraction (operating cost). Optimum combination of flow rates, number of stages and degree of solute transfer is decided based on economics including cost of solvent recovery. The theoretical number of stages, N, required for separation is estimated by graphical construction on triangular diagram or rectangular ternary plot as per convenience. As discussed in Chapter 9, the ternary rectangular plot is commonly used with the vertical and horizontal axes denoting the respective equilibrium concentration of solute B and extraction solvent S. The composition in the raffinate and extract phase is denoted in Fig. 13.5 by x and y, respectively, and the subscripts denote the components. The computation is based on locating the mixture point M and the difference point D. In the most general case, all three components (A, B and S) may be present in the feed as well as the solvent used for extraction. Presence of A and B in the solvent stream is particularly common as the recovered solvent, often reused, may still contain small amount of A and B. Design input

Stream

Flow rate (kg/hr or kg mol/hr)

% A

% B

% S

Feed

F

zA;F

zB;F

zS;F

Extraction solvent

S or ENþ1

yA;S

yB;S

yS;S

Raffinate

R

xA;R

xB;R

xS;R

Steps to find the number of stages N 1. Desired composition of raffinate is usually specified in terms of concentration of B in solventfree basis and so this composition lies on the y-axis and is located at R0 . 2. The feed and the solvent point (F and ENþ1 ) are located on the diagram (Fig. 13.5A).

13.4 Design of continuous countercurrent contactors

399

3. The mixture point M is located from the flow rates ðF and SÞ and composition of feed and solvent as shown below. Ordinate: zB;M ¼

F  xB;F þ S  yB;S FþS

(13.2a)

zS;M ¼

F  xS;F þ S  yS;S FþS

(13.2b)

Abscissa:

In case pure solvent is used and the feed stream is free of solvent, yS;S ¼ 1 and xS;F ¼ 0 4. The line joining ENþ1 and R0 intersects the raffinate curve at RN . 5. The tie line that passes through point F (extended beyond the envelope) is drawn and its extract end is E1min . Coordinates of E1min denote extract composition when minimum amount of solvent stream is used to produce raffinate of desired composition ðR0 Þ. 6. The line joining ENþ1 and F intersects the line RN E1min at Mmin which corresponds to the mixture composition using minimum amount of solvent stream. ðFM

Þ length

min 7. Minimum solvent to feed ratio is calculated as rmin ¼ ðENþ1 M , based on the ‘lever arm min Þ length rule’. 8. Operating solvent to feed ratio is calculated as r ¼ rmin  k, where typically 1.1 < k < 2. The value of r should be less than rmax which corresponds to maximum solvent required. This is obtained when M lies on the extract curve. Exceeding this upper limit on solvent shall lead to a mixture that shall not separate as extract an raffinate phases. 9. Based on r, composition of mixture M is located as ðFMÞlength ¼ r  ðENþ1 FÞ length. One may note that coordinates of M can also be obtained from expressions based on balance of component B and S. 10. The line RNM is drawn and extended to meet the extract curve at E1 that corresponds to the operating extract composition from the process. [The construction may be continued on the same graph for estimating the number of stages (N), but to avoid clumsiness a separate graph as in Fig. 13.5B is often used to locate the difference point ðDÞ and continue with the construction.] 11. On the new graph, all except RN E1min line and the tie line through E1min is drawn. 12. Lines E1 F and ENþ1 RN are extended to meet at the operating difference point ðDÞ as shown in Fig. 13.5B. 13. The tie line ðE1 R1 Þ through point E1 is drawn to obtain raffinate composition leaving stage 1 denoted as R1 . 14. In order to locate details of the next stage, a line through the point D and R1 is extended to meet extract curve at E2 . 15. The tie line ðE2 R2 Þ through point E2 is drawn to obtain raffinate composition leaving stage 2 denoted by R2. Tie lines are conveniently drawn with the help of an xey plot (denoting equilibrium composition of solute B in the raffinate and extract phase) drawn beside the ternary rectangular plot. The procedure is illustrated in Figs. P13.2 and P13.3 and outlined below. A horizontal line is drawn from any point, say E on the extract curve to intersect the diagonal of the xey plot. This locates yE ¼ xE . A vertical line downward from yE ¼ xE point intersecting the equilibrium curve locates xE in equilibrium with yE . A horizontal line from point xE to the

400

Chapter 13 Extraction

B

(A)

100% m Min. .. solvent = F ×

FMmin length EN+1 Mmin length

m Actual solvent = k × Min. .. solvent

xB, yB →

FM length Extract =F× EN+1 M length curve Raffinate E1min curve P Tie line F

E1 Mmin

R'

RN

M

EN+1 S 100%

0% A 0% xS, yS →

B

(B)

100%

feed

=

xB, yB →

FM length

solvent

Extract curve Raffinate curve P

EN+1 M length

Δ xS,Δ = xB,Δ =

(F.xS,F – E1.yS,E1) (F – E1) (F.xB,F – E1.yB,E1) (F – E1)

R'

,abscissa

,ordinate

Tie line

F

0% A 0%

R1 RN R2

E1 E2

M Tie line

EN+1 S 100%

x S, y S →

FIGURE 13.5 Countercurrent multistage extraction (A) locating Mmin (B) estimating N.

raffinate curve locates xE in the ternary rectangular plot. Line yE xE drawn from the extract to the raffinate curve is then the required tie line through point E. 16. Steps 14 and 15 are repeated till stage N, where the composition of raffinate falls below RN . In the case shown in Fig. 13.5B, the condition is satisfied for N ¼ 2.

13.4 Design of continuous countercurrent contactors

401

17. The composition of extract and raffinate streams can then be read off the graph. Often the number of stages is estimated by available shortcut methods. One such method is the McCabeeThiele construction of stages discussed in Chapter 11 for straight operating and curved equilibrium line. The technique can be used here for constant flow rate of feed solvent F 0 and extract solvent S0 where the solute concentrations are given as weight ratio (X) of solute to feed solvent in raffinate and ratio (Y) of solute to extraction solvent in extract. For perfectly immiscible solvents, both flow rates F 0 and S0 remain constant in all the stages and a linear operating curve of slope ðF 0 =S0 Þ ¼ ðE0 =R0 Þ is obtained. Material balance from the feed end to stage n gives Ynþ1 ¼

F0 E0 YE  F 0 XF X þ n S0 S0

(13.3a)

and from raffinate end to stage n gives F0 E 0 Y E  R0 XR X þ (13.3b) n1 S0 S0 Thus, from overall material balance, the end points of the operating lines XR YS and XF ; YE are related as Yn ¼

F 0 XF þ S0 YS  R0 XR (13.3c) E0 The equilibrium curve is plotted in the XeY coordinates and the number of stages is obtained by McCabeeThiele stepwise construction discussed in Chapter 10. An example plot is shown in Fig. 13.6. YE ¼

1.0

(XB,F,YB,E)

Y

1

2

(XB,RN,YB,RN+1)

0 1.0

0 X

FIGURE 13.6 McCabeeThiele construction for finding number of stages.

402

Chapter 13 Extraction

For a linear operating and a linear equilibrium curve of slope m passing through the origin, the number of theoretical stages can be estimated from Kremser equation as function of the extraction factor ε which is the ratio of the slope of the equilibrium line and the operating line, i.e. ε ¼ mS 0 =F 0 . For ε s 1,      XF  YS =m 1 1 ln þ  1 XR  YS =m ε ε (13.4a) N¼ ln ε For ε ¼ 1, N¼

XF  YS =m 1 XR  YS =m

(13.4b)

For non-zero intercept of equilibrium line; m0 the partition ratio in coordinates XeY should be used instead of m. For nonlinear equilibrium curve, the geometric mean of m at the concentration leaving pffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi stage 1 and stage N is used, i.e., m ¼ m1  mN . Calculation of differential countercurrent extraction (packed column) is done by NTU method discussed in Chapter 10. This is adopted if data on mass transfer coefficient is available. The number of transfer units can be defined based on extract ðNtoE Þ or raffinate phase ðNtoR Þ as Z x0 dx NtoR ¼ (13.5a) x  x xN Z y0 dy (13.5b) NtoE ¼ y y yN where x is fraction (weight/mole) solute in raffinate phase and x is fraction solute in raffinate phase in equilibrium with extract phase of composition y. For a linear operating and equilibrium curve with zero intercept, NtoR can be obtained from Kremser equations (Eq. 13.4). The denominator of Eq. (13.4a) is only different for εs1:0.     XF  YS =m 1 1 ln þ 1 XR  YS =m ε ε NtoR ¼ (13.6) ð1  1=εÞ The difference is more pronounced for high ε where  ε  NtoR ¼ N  ln 1ε This gives the column height as

(13.7)

L ¼ HtoR  NtoR ¼ HtoE  NtoE

(13.8a)

L ¼ NTS  HETP

(13.8b)

and in terms of NTS as Presence of significant axial dispersion in large diameter towers of actual plant decreases the effective concentration driving force and so the tower height to be provided is larger than that predicted by plug flow NTU. This is incorporated by adding height of dispersion units HDU to transfer units, viz

13.4 Design of continuous countercurrent contactors

ðHTUÞov ¼ HTU þ ðHDUÞc þ ðHDUÞd

403

(13.9)

where subscripts c, d and ov denote continuous phase, dispersed phase and overall, respectively. Staged equipment are also best modelled by NTU method for low-stage efficiency when extraction factor ε >> 1:5. The overall stage efficiency is defined similar to staged distillation and absorption towers, as the NTS to the number of actual stages and expressed as %. Commonly used extraction equipment is designed to achieve the equivalent of 1e8 theoretical countercurrent stages. The overall efficiency of mixer settler is more than 80% (usually around 90%e95%) and that of sieve tray is typically 8%e30%.

13.4.2 Design parameters for extraction towers Typical design parameters of extraction towers are discussed in the subsections below. Packed tower Unlike gaseliquid and vapoureliquid contactors, the interfacial area for mass transfer in packed extraction tower is not the packing-specific surface area. Rather, the mass transfer area is almost independent of packing surface area and is a function of dispersed phase holdup. Packing in this case merely increases the turbulence and also serves to redisperse the coalesced droplets. Both random and structured packings can be used to cause droplet coalescence and breakup. Structured packings provide higher values of turndown (operational flexibility), allowable bed height and throughput. The throughput is 40e80 m3/m2 hr while that of random packed towers is 20e30 m3/m2 Packing characteristics hr. Nevertheless, random packing is still a choice in industries. This is mainly because it is cheaper and easy to clean. The HETP is typically 0.4e1.5 m for random packings (overall height of transfer unit HOL: 0.9e1.7 m) and 0.5e1.6 m for structured packing. Raschig rings were the most common random packing but now Berl saddles are popular. In LLE. corrosion and wetting are two major considerations for selection of packing material. The packing needs to resist corrosion and be well wetted by the continuous phase. Strength is also a consideration and ceramics is usually avoided. This is because ceramics is brittle and often breaks and clogs the flow and pump suction strainers and valves. Nevertheless, they are used for corrosive liquids when aqueous phase is continuous. Stoneware packings and metal packings are used when water is the continuous phase and carbon rings or saddles are adopted for continuous organic phase (say toluene). Similar to absorption towers, the ratio of tower diameter to packing size should be greater than 8:1 to minimise wall effect. Pilot columns usually use smaller packings, or else an unnecessarily large diameter pilot column is required. In case structured packing is selected, the test columns can be equipped with industrial packing. This minimises scale-up risk on throughput but limits the minimum pilot column diameter to 50 mm and also requires efforts to minimise wall bypassing. Liquideliquid flooding velocity in packed bed can be predicted from the flooding velocity correlation plot in packed extraction towers correlation (McCabe and Smith, Unit operations of Chemical Engineering, McGraw Hill Int’l Student’s Edn., 4th edition, pg 542, 1985), an analytical form of L-L flooding in packed bed which is

404

Chapter 13 Extraction

C ¼ exp 8.8082  0.0563  flnðDÞg2  0.4981  lnðDÞ  0.2     s mc a 1.5 where D ¼   rc ε Dr

(13.10) (13.11a)

and C¼

pffiffiffiffiffiffiffiffi pffiffiffiffiffiffiffiffi 2 Us;d þ Us;c rc amc

(13.11b)

The above is valid for 1 < C < 1000. It is important to note that the group D is not dimensionless and the following units are to be used. Us;d ; Us;c the superficial velocity (volumetric flow rate per unit tower cross section) of dispersed and continuous phase in ft/hr, rc is density in lb/ft3 and mc is viscosity of the continuous phase in lbm/ft hr. a the specific surface area of packing is in ft2/ft3 and ε is the packing void fraction. The values of a and ε provided in Chapter 14 for gaseliquid packed towers can be used for extractors as well. Dr ¼ rc wrd and s is interfacial tension in dyne/cm. While working with SI units, the following conversion factors may be used: 1 m/sec ¼ 11,811 ft/h 1 kg/m3 ¼ 0.062428 lb/ft3 1 kg/(m sec) ¼ 2419.08833 lb/(ft hr) and 1 cP ¼ 2.419 lbm/(ft hr) 1 m2/m3 ¼ 0.3048 ft2/ft3 Most critical data for an accurate estimation of flooding capacity is the value of interfacial tension that governs the droplet size. In absence of experimental data, s in dyne/cm can be estimated for a ternary system of solvent A, solvent S and solute B as   ðxBA þ xBS Þ s ¼  7.34 ln xAS þ xSA þ  4.90 (13.12) 2 where xAS is the mole fraction of solvent A in saturated solvent-rich S, xSA is the mole fraction of solvent S in saturated solvent-rich A and xBA and xBS are the mole fractions of solute B in A and S, respectively. Eq. 13.12 is applicable for 4  s  52:5 dyne/cm and predicts s within an uncertainty around 15% for ternary systems with partially miscible solvents and solute completely miscible in both solvents. The design flow rate is recommended at 50% flooding or lower, even up to 20% in unknown systems. Packed columns with D < 800 mm are usually flanged from cost considerations, while in higher diameter columns, the sections are welded. Manholes are provided for inserting the internals in pieces inside the welded columns. Nozzles for feed, draw off, level measurement and temperature control are required. Often inspection glasses are provided but these require cleaning due to accumuColumn accessories lation of dirt settling from the liquids. Tapping nozzles at various elevations are provided for sample withdrawal from the column. These are called tryline. Samples from the try-lines are visually compared to identify the interphase location between two consecutive try-lines. This is important for systems that are dark, dirty or when inspection glasses cannot be provided usually for safety reasons.

13.4 Design of continuous countercurrent contactors

405

Performance of a column strongly depends on the quality of flow distribution. Large diameter columns should have sufficient pressure drop to ensure good radial distribution at turndown conditions while avoiding formation of spray caused by high velocity through nozzles at design rates. Sufficient open area is also necessary. Flow distributor Similar to vapoureliquid and gaseliquid distribution, redistributors are required after 1.5e3 m to ensure proper drop size distribution. Long bed heights (>10e12 m) with structured packing are not advisable. Sieve tray tower These are similar to the trays used for distillation and absorption. The lighter liquid is dispersed as droplets while passing through the sieve trays. They coalesce and jet above each tray and get redispersed as they rise through the sieves of the upper tray. The heavy liquid flows as the continuous phase across the tray. The tray spacing (TS) is 10e25 cm and the perforations on the tray (do) are 3e7 mm in diameter. Tray active area (Aa) is 15%e25% of column cross-sectional area. Typically, downcomer area (Ad) is 5%e10% of column cross-sectional area. Overflow weirs are not used on downcomers in liquideliquid service. Fractional tray efficiency h can be estimated from pffiffiffiffiffiffi 0.2945 TS h¼ ðUd =Uc Þ0.42 (13.13) sd00.35 where s is in dyne/cm, TS and do are both in m and Ud and Uc are in m/sec. Typical tray efficiency is 8%e30% and HETP is 0.8e1.2 m. Systems with high interfacial tension require larger HTU and the stage efficiency is low. Sieve plate towers have capacities ðUd þ Uc Þ of 27e60 m3/(m2 hr) and are scaled-up keeping the mixture superficial velocity ðUd þ Uc Þ same. Flooding occurs when either the liquideliquid interface does not remain at the tower bottom or a large portion of dispersed layer accumulates on the trays and restricts the flow of the continuous phase. Proper tray design solves the dispersed-phase accumulation problem. To keep the interphase close to the tower bottom, the dispersed-phase flow rate should be high enough to overcome the head losses due to (a) Flow of dispersed phase through perforations on each plate, ho (obtained from standard orifice equation). ho ¼

rd Uo2 ð1  A0 =AÞ2 2gDrCo2

(13.14)

Typically Uo is in the range 15e30 cm/s and Co w0:67. (b) Interfacial tension hs , important when continuous phase preferentially wets the tray. hs ¼

4s Drdo

(13.15)

406

Chapter 13 Extraction

(c) Flow of continuous phase hc which includes friction on plate, in downcomer (negligible), contraction and expansion at downcomer entry and exit and abrupt changes in direction. This is considered to be 4.5 velocity heads. hc ¼

4.5  Ud2  rc 2g  Dr

(13.16)

In Eq. 13.14e13.16, the terms with subscript ‘o’ refer to parameters at the perforations and Co is the discharge coefficient. The depth of liquid layer for preferential wetting of tray by continuous phase is the higher of hs and ðho þ hc Þ. It is important check that TS > ðhc þ hd Þ, where hd ¼ ho þ hs . This avoids excessive entrainment and break up of dispersed-phase stream into droplets before entering the next plate. Spray extraction tower The following are the typical features: -

One or two stages Large axial dispersion in continuous phase Range of ðUd þ Uc Þ is 15e75 m3/(m2$hr) HOL and HETP: 3e6 m

Agitated tower Typical design specifications of pulsed towers are listed in Table 13.3 and that for Scheibel tower, Karr reciprocating plate tower and rotary disc contactor are mentioned in Table 13.4. Schematic diagrams of pulsed tower, Scheibel tower and rotary disc contactor with typical internals and key dimensions marked are shown in Fig. 13.3AeC, respectively.

Table 13.3 Typical design specifications of pulsed sieve tray and packed towers.

Equipment

Pulsation Specification

Tower details

ðUd DUc Þ m3/(m2 hr)

Flooding velocity ðUd DUc ÞF m3/(m2 hr)

HETP (m)

Pulsed sieve tray tower

6e25 mm amplitude 100e250 cycles per minute

TS 50 mm, Aa 0.2e0.25 A do 3 mm No downcomer

25 to 35

60

0.15e0.3 m

Pulsed packed tower

6e8 mm amplitude Frequency Dtank . Settling and coalescence may occur in the dispersion tank after agitation has stopped or the two-phase mixture can be directed to a gravity settler after the dispersion section. Both

408

Chapter 13 Extraction

types can be used in series for countercurrent operation. Such a countercurrent arrangement is used in extraction of uranium or copper salts from aqueous solution. Gravity settlers are vertical/horizontal tanks where the liquideliquid dispersion is continuously settled and coalesced and the settled liquids are continuously withdrawn. The inlet velocity must be kept low for minimum interfacial disturbances. (See Chapter 17 for sizing etc.) Impeller: Mixing in the tank is by agitator/impeller which may be closed or open type. It is mounted on a shaft and driven by an electric motor. A properly designed impeller should impart circulatory motion in both radial and axial direction to prevent gravity settling and separation due to centrifugal force (as heavier liquid has a tendency to form a layer close to the tank wall). While adequate dispersion is necessary for rapid extraction, it is important to prevent formation of stable emulsions. Impellers are usually classified based on the type of flow they induce e tangential (paddle), axial (propeller) and radial (turbine). The typical characteristics of propeller and turbine are listed in Table 13.6 and their schematics are presented in Figs. 13.7 and 13.8. Table 13.6 Typical characteristics of propeller and turbine type impeller. Propeller

Turbine

· Most economical for low viscous liquids in small tanks · Suitable ( tank diameter D, twoethree impellers on same shaft with bottommost impeller located DeD/3 above the tank floor High speed

Turbine: The geometric parameters defining the turbine agitator are shown in Fig. 13.7 and their typical standard dimensions are as follows with 4 baffles. DI WI 1 1 Ddisc 2 L1 ¼ 0:3  0:5; ¼ to ; ¼ ; ¼ 0:25 5 3 DI D DI 8 DI H H 1 1 Wb 1 1 ¼ 1; ; tgap ¼ 0:1  0:15Wb ; Nb ¼ 4 ¼ ; ¼ to D 3 D 8 12 D

13.5 Design of mixer-settler

409

tgap

Wb

Baffle LI H WI Ddisc HI

DI

D

FIGURE 13.7 Standard six flat blade turbine agitator with disc.

Baffle

H≈D

HI

DI D

FIGURE 13.8 Schematic of three-blade propeller agitator.

Propeller: Marine type propellers are more common. These have square pitch, i.e., in one revolution, the propeller pushes the liquid forward by a distance P (Pitch) equal to its diameter DI . Usually DI DI 1 1 D  5 is adopted for single-phase mixing but for two phases D ¼ 3 is used. The preferred propeller elevation ðHI Þ above the vessel floor is DI , though the propeller may also be submerged by a depth DI from the top liquid surface. Paddle type is generally not recommended for extraction as they have low speed (20e200 rpm) and produce poor circulation. These are used for viscous liquids like starch, paste, paints, adhesives,

410

Chapter 13 Extraction

cosmetics, etc. Modified paddles, e.g., anchor agitator, can mix viscous liquids (50,000 to 500,000 cP). For still higher viscosities (usually 500,000 to 1,000,000 cP and up to 25,000,000 cP), helical ribbon agitators (ribbon formed in a helical path and attached to a central shaft), which have low rotation speed, are used. Baffles Vortices are inevitable with any impeller type except for highly viscous liquids or marine propellers in an off-centre arrangement. These are eliminated by installation of baffles that are narrow, flat strips welded or fastened vertically along the tank wall. Typically, four equally spaced vertical baffles of length equal to liquid depth and width equal to D=10 to D=12 are arranged radially around the vessel diameter. This arrangement has been observed to produce negligible swirl and no further advantage is obtained with additional number of baffles.

13.5.1 Holding time The mixer settler can be sized from an estimate of holding time obtained from mass balance equation for mixing and from terminal velocity of droplets for settling as outlined in Chapter 17. For N moles of solute transfer from raffinate to extract (denoted by subscripts R and E, respectively) during time t, if phases attain equilibrium, the concentration of solute changes from coE to cE in extract and coR to cR in raffinate. Considering a linear equilibrium relationship within the range of concentration change, cE ¼ mcR

(13.17)

Assuming immiscible extract and raffinate phases with flow rates E and R, respectively,    c E cE  c0E ¼ R c0R  E (13.18) m  dN dcE ¼E ¼ KE a cE  cE dt dt

(13.19)

K E a is the overall volumetric mass transfer coefficient based on overall concentration gradient where cE cE in the extract phase as discussed in Chapter 10. This gives E  t¼ lnð1  hÞ mE ðKE aÞ 1þ R

(13.20)

where N is the moles solute transferred in the stage in time t and N  is the moles solute transferred till  cE  c0E N . equilibrium is reached. This gives stage efficiency as, h ¼ N  ¼   cE  c0E Eq. 13.20 can also be obtained in terms of a raffinate rate coefficient as   R  1þ R mE t¼  lnð1  hÞ (13.21) ðKR aÞ

13.5 Design of mixer-settler

411

The holding time in mixer is usually 1e3 min except in cases like metal extraction where due to reactive extraction, a holding time of 10e15 min is usual. It is important to consider settling along with mixing time. Often a higher intensity of mixing reduces residence time for mass transfer while creating fine droplets difficult to separate. Droplet size is a design parameter to estimate settling time. Small droplets achieve large interfacial area and faster extraction while requiring a longer settling time. Most stable emulsions are characterised by maximum droplet diameter of the order of 1e1.5 microns while diameters around 1 mm produce relatively coarse dispersions that separate readily. Coarse dispersions are preferred in extraction and reacting systems involving two liquid phases that need to be separated.

13.5.2 Power and mixing time Agitator design needs to consider power requirement as well as adequacy of mixing since higher power consumption is not necessarily associated with adequate mixing. Several empirical correlations are available to predict the power required. A typical correlation for commonly used turbines and propellers (Fig. 13.9) with Newtonian liquids in baffled, cylindrical vessels correlates the dimensionless     D2 N r Power number, NP ¼ D5 NP3 r with impeller Reynolds number, ReI ¼ I m I m for axial impeller I

I

m

m

shaft and liquid depth equal to tank diameter. The parameters defining the numbers are P (power in W), DI (impeller diameter in m), NI (impeller rotational speed in revolutions per second), rm (mixture density in kg/m3), and mm (mixture viscosity in kg/m sec). The flow is laminar for ReI < 10 and turbulent for ReI > 104. In the range 10 < ReI < 104 , the flow is transitional, i.e., turbulent at the impeller and laminar further away. For immiscible liquid mixtures, rm and mm can be expressed in terms of individual liquid properties and mixture composition expressed as volume fraction vE and vR of the extract and raffinate phases. Note that vE þ vR ¼ 1.

For baffled vessels mm ¼

rm ¼ vE rE þ vR rR

(13.22)

mm ¼ mvEE mvRR

(13.23)

  mc 1.5md a 1þ mc þ md ð1  aÞ

(13.24)

where a is the dispersed-phase fractional holdup in the vessel. Viscosity of individual phases with n components with xi mole fraction of each component, may be n P found from the mixing rule: lnðmÞ ¼ ðxi  lnðmi ÞÞ. 1

For high agitation speed, emulsion formation may increase mixture viscosity and Eq. 13.24 does not give accurate results. Typically, P ¼ 0:8 to 2 kW/m3 of fluid for intense agitation required in extraction where P is the power imparted to liquid by impeller and does not include losses in the motor, speed reducing gears, bearings, stuffing box, etc., which is around 0.3 to 0.4 times P. The power required for liquids of low to moderate viscosity is 0.1e0.2 kW/m3 for mild agitation and blending and 0.4e0.6 kW/m3 for vigorous agitation. For flat six-blade open turbine, deviations from the standard design influences Power number as follows: (a) NP fWI =DI

412

Chapter 13 Extraction

(b) NP almost same for DI =D in the range 0.25e0.5 (c) For two turbines installed on same shaft with spacing w DI, the total power is 1.9 times the power required for single impeller where spacing refers to the vertical distance between the  bottom edges of the two turbines. This also applies to six-blade pitched blade (45 ) turbine.

P

Np =

5 3 DI .NI . ρm

102

101 1 2 3 4

100

5 10–1 100

101

102 ReI =

103

104

105

106

(DI2.NI. ρm/μm)

FIGURE 13.9 Correlation of NP versus ReI for different impeller types in baffled agitators. (1) Flat six-blade turbine with disc DI =WI ¼ 5; 4 baffles with ðD=Wb Þ ¼ 12, (2) flat six-blade open turbine with disc DI =WI ¼ 8; 4 baffles with ðD=Wb Þ ¼ 12, (3) six-blade turbine (45 blade), DI =WI ¼ 8; 4 baffles with ðD=Wb Þ ¼ 12, (4) propeller turbine with Pitch ¼ 2  DI ; 4 baffles with Pitch ðD=Wb Þ ¼ 10, also valid for off centre in angular off-centre position without baffles, (5) propeller turbine Pitch ¼ DI ; 4 baffles with ðD=Wb Þ ¼ 10, also valid for off centre in angular off-centre position without baffles. From sources - Curves 1,2 & 3 are from Bates, R.L., Fondy, P.L. & Corpstein R.R. (1963). Examination of some geometric parameters of impeller power. Industrial & Engineering Chemistry Process Design and Development. 2(4), 310e314. Curves 4 and 5 from Rushton, J.H., Costich, E.W. & Everett, H.J. Power characteristics of mixing impellers. Chemical Engineering Progress. 46, 395e404, 1950; 46:467e79, 1950. [Mixing Equipment Co., Rochester, NY, and Illinois Inst. Technology, Chicago, IL].

Same NP read from Fig. 13.9 can be used for Re < 300 in case of unbaffled vertical and horizontal cylindrical vessels and for baffled square section tanks. The power consumption in unbaffled vessels is less as compared to baffled vessels when Re > 300. For helical ribbon agitator, the NP  ReI correlation is as follows for ReI < 20(viscous liquids) NP ¼

186 ; for agitator pitch ¼ D ReI

(13.25)

NP ¼

290 D ; for agitator pitch ¼ ReI 2

(13.26)

and the typical dimensions are

DI D

¼ 0:95 with minimum 0.75 and

WI DI

¼ 0:095.

13.5 Design of mixer-settler

413

The efficacy of mixing expressed as dimensionless mixing factor fmix can also be correlated with ReI for a turbine agitator in a baffled tank where  2 1 1 NI D2I 3 g6 D2I fmix ¼ tmix (13.27) 1 3 H 2 D2 and tmix is the mixing time in seconds. The graphical correlation presented in Fig. 13.10 shows that fmix 2

and tmix NI3 is constant for ReI > 1000.

3/2

.D

102

H

1/2

fmix = tmix ×

(NI.DI2)

2/3

.g

1/6

.DI

1/2

103

101

100 100

101

102

103

104

105

106

ReI = DI2.NI. ρm/μm

FIGURE 13.10 Mixing time as function of Reynolds number for miscible liquids in a baffled vessel agitated by turbine impeller. From Norwood, K.W., & Metzner, A.B. (1960). Flow patterns and mixing rates in agitated vessels. AIChE Journal, 6, 432.

13.5.3 Scale-up For most LLE applications, the mixing section is scaled based on geometric similarity, i.e., in pro1

portion to increase of vessel diameter for D ¼ H1 or as (vessel volume) 3 with the same power per unit volume and equal holding time. It is difficult to achieve dynamic (ratio of viscous, inertial or gravitational forces) and kinematic (velocity ratio) similarity at the same time. 2 Using the same scaling for all dimensions shown in Fig. 13.7, DI scales as D while NI fð1=DÞ3 for equal mass transfer rate or equal power per unit volume. For fully developed turbulence, PfNI3 and 11

tmix can be scaled for similar geometry and same power per unit volume in turbulent flow as tmix f DI8 , 11

while for same tmix ; VPfDI4 . Usually agitators are scaled at constant power per unit volume which gives a larger mixing time for larger vessels. Scale-up to continuous process can also be performed from approach to equilibrium in batch ktb , and approach to equilibrium in continuous process E o process Eb where Eb ¼ ccii c cont is c ¼ 1  e ktcont Econt ¼ 1þktcont .

414

Chapter 13 Extraction

The value of k as obtained from batch data (mixing time tb in batch process) is used to compute residence time tcont for continuous process. From residence time and power input, scale-up is done based on geometric similarity and equal power input per unit volume.

13.5.4 Flow mixers These are widely used for refining light petroleum distillates with small quantities of solvent. Dispersion is achieved by ‘in-line’ or flow mixers where agitation occurs by fluid flow. Their use is limited to low viscosity liquids ( 10. Design of new systems without pilot data can be made considering that flow capacity Ud þ Uc and HETS increases with decrease in N and Dr and increase in Ds and H1 . HETS passes through a minimum with increase in N.

Further reading

425

Further reading Wankat, P. C. (2006). Separation process engineering (3rd ed.). Pearson Education. Treybal, R. E. (1963). Liquid extraction (2nd. Ed). McGraw-Hill Classic Textbook Reissue Series. Koncsag, C. I., & Barbulescu, A. (2011). Liquid-liquid extraction with and without a chemical reaction. In Mass transfer in multiphase systems and its applications. IntechOpen. Perry, R. H., Chilton, C. H., & Kirkpatrick, S. D. (Eds.). (1963). Chemical engineers’ handbook (4th ed.). New York: McGraw-Hill. Geankoplis, C. J. (2003). Transport processes and separation process principles:(includes unit operations). Prentice Hall Professional Technical Reference.

CHAPTER

Column and column internals for gaseliquid and vapoureliquid contacting

14

14.1 Introduction The terms columns and towers are synonymous. These are continuous interphase mass transfer equipment introduced in Chapter 8 (Interphase Mass Transfer, Table 8.1). As mentioned, towers for vapoureliquid and gaseliquid contacting can be spray towers, bubble columns, packed towers and tray towers. Spray columns and bubble columns in contrast to the rest do not have any special contacting internals. Their fluid inlet/exit arrangements and distribution is common with the packed and tray towers. This chapter focuses on the tray and packed towers e the most common industrial equipment for gaseliquid contacting. Tower design aims to enhance effective contact between the process streams and reduce entrainment while minimising pressure drop. In addition, the design must be consistent with the economics dictated by the process and type of construction. Tray internals promote formation of vapour bubbles, and mass transfer occurs across the bubble surface in a liquid pool on the tray. Some internals are used to improve physical separation as well as other miscellaneous purposes including routing / diversion of phases. Schematic of a tray tower typical for fractionation service is shown in Fig. 14.1A. Trays in a tower are perforated and the perforations may or may not be fitted with attachments. This treatise focuses on the basic types of vapoureliquid and gaseliquid tray towers namely sieve tray, bubble cap and valve tray. Fig. 14.2 shows the tray construction features namely tray support ring (TSR) welded to the shell and major and minor beams supporting the tray from bottom. The tray is built in suitable sections, one of which acts as a manway. • •

•

Sieve trays have simple circular perforations on the tray deck. Bubble cap trays have caps covering individual vapour riser tubes fixed to the tray deck. The assembly is submerged in the liquid pool. Vapour escapes as a chain of bubbles through slots on the side of the caps. Valve trays have flat or slightly curved plate concave to the deck as cover on the perforations in the tray deck. The valve plate has three legs that pass through the hole and are bent outwards. These restrain the plate from being lifted beyond a maximum limit by vapour flow through the hole. The valve plate edge is also provided with three crimps that do not allow it to flush with the deck and a small clearance (w2e3 mm) remains, through which the tray deck liquid may drain when the tower is shutdown. Similar to the aforementioned trays, vapour escapes as a train of bubbles through the liquid pool.

Process Equipment and Plant Design. https://doi.org/10.1016/B978-0-12-814885-3.00014-2 Copyright © 2020 Elsevier Inc. All rights reserved.

427

428

Chapter 14 Column and column internals for gas

Packed towers house a bed of packing irrigated by downflow of liquid and the vapour flows through the rest of the void space in the bed. Small packing elements dumped randomly in the tower constitute a ‘dumped bed’ of ‘random packing’. Towers can also have packing elements stacked regularly that makes a ‘stacked packing’ section. A packed bed can also be blocks of ‘structured packing’ usually sewn together with stainless steel wire and snugly fitting the tower cross section. Grid packing is another type that has lower pressure drop, higher capacity and much lower efficiency than the other types. Fig. 14.1B depicts a packed column.

(A)

(B) Overhead

Overhead vapour

vapour

Demister Reflux

Distributor

Reflux

Rectifying section

Tray Packed bed (random/structured)

Feed

Feed

Stripping section

Bed limiter

Tray Packing support

Reboiled vapour

Reboiled vapour Bottoms

Bottoms

FIGURE 14.1 Typical distillation column configuration for (A) tray and (B) packed column.

14.1 Introduction

429

Manway

Downcomer and weir Calming area Major beam Tray support ring

Major beam clamp, welded to tower wall Major beam Minor beam support clamp Subsupport plate ring used with angle ring

Minor beam support clamp Peripheral ring clamps

Minor beam support clamp

Subsupport angle ring

FIGURE 14.2 Typical tray construction.

Considerations for choosing between packed and tray towers: (i) Pressure drop per theoretical stage of contact for the same vapour and liquid load: pressure drop increases in the order Structured packing (w15 mm WC) > Random packing (w30 mm WC) > Sieve tray (w50e100 mm WC)> Valve tray > Bubble cap tray. (ii) Capacity: Lower pressure drop can be translated into higher throughput or capacity of the tower and revamp for capacity augmentation of existing tray towers with some of the trays replaced by packed sections of structured packing is common. (iii) Energy consumption: Lower pressure drop in packed towers allows the tower bottom pressure to be lower. This translates into higher volatility difference between key components in the lower part of the tower, thus requiring lower bottom temperature, reflux flow, condenser and reboiler heat loads, all of these lead to lower energy input. (iv) Maintenance and inspection: Towers are drained prior to maintenance shutdown. Complete draining of packing is more difficult as compared to trays. Any hazard due to the presence of the remaining liquid needs to be handled carefully by washing with utility water/steaming. (v) Material of construction: Often packing material is cheaper compared to trays. This is particularly relevant for corrosive services, e.g., ceramic packing elements are resistant to corrosion by acid and alkali. (vi) Liquid holdup: Packed tower has lower holdup than tray tower and is, therefore, a preferred choice for (i) batch distillation to produce higher yield of distillate and (ii) thermally unstable substances as the lower residence time deters its decomposition. On the other hand, for cases where higher residence time is desirable, e.g., chemical reactions, absorption, etc., tray towers

430

(vii) (viii) (ix)

(x)

(xi)

(xii) (xiii) (xiv)

(xv)

(xvi)

Chapter 14 Column and column internals for gas

are favoured. High liquid holdup on tray as well as downcomer also dampens composition fluctuations. The liquid holdup and residence time on a tray are primarily decided by height of the outlet weir. Change in feed composition: Corrective action for changes in feed composition can be made more easily in tray towers by altering the reflux ratio and/or relocating the feed stage. Foaming and emulsion formation tendency: This is higher on trays due to higher vapour velocity while passing through tray perforations/slots. Tower diameter: Tray towers are preferred for large diameter, while packing is a suitable option for low throughput. It is difficult to inspect/maintain tray columns smaller than 1000 mm diameter due to limited space. In such cases, either an oversized tower or tray cartridges fitted inside the column are suitable options. The latter option often does not have perfect sealing at the wall that may lead to lower tray efficiency. Large diameter e packed towers often suffer from maldistribution of vapour and liquid flow leading to lower efficiency compared to tray towers. Large diameter tray design needs to ensure that the difference in liquid depth across the flow path (on cross-flow trays) is within limit to minimise channelling of vapour from areas with lower liquid depth. Multipass trays and cascade trays are design options to reduce the difference in liquid depth. Temperature cycles: Cyclic and large temperature variation from cold start-up condition to high operating temperatures damage (squeeze/crush) packing due to differential expansion and contraction over the thermal cycles. Packed towers may not be suitable for such services. Presence of solid: Systems with solids are better handled in tray towers than in packed columns as the velocities are higher. Higher velocity and agitation tends to keep the particles suspended rather than allowing them to settle down and clog the system. Cleaning of trays is much easier. Solids often clog the liquid distributor in packed towers. Heat removal capacity: Coils can be incorporated on trays to allow heat addition or removal. Weight: Usually tray towers weigh less than packed towers. This may reduce foundation cost. Liquid flow rate: High liquid rates in packed section reduce the contact efficiency, particularly in structured packing. With options of cascade and multipass trays, tray towers can handle higher liquid rates more efficiently. Turndown (or turndown ratio): There are two definitions of ‘Turndown’ in literature and one is the inverse of the other. In this book, turndown is defined as the ratio of normal operating (intended) vapour throughput to the minimum allowable vapour flow rate without significant loss of tray efficiency. Expectedly, this ratio is above one. A design with a higher ‘turndown’ suggests feasibility of operating the tower (without significant loss in separation efficiency) up to a lower throughput limit. The turndown range of tray towers is higher. Packed towers do not perform satisfactorily at lower liquid rates. Typically below 60% of design capacity, their performance gets affected primarily due to improper distribution of liquid from the distributor. Flexibility in operation: Tray towers can handle liquid load over a wide range, while gas load must lie within a relatively narrow range for achieving the design tray efficiency, only valve trays allow greater operational flexibility. On the other hand, packed towers are extremely flexible as far as gas load is concerned but require a minimum liquid load. This often excludes their use in vacuum operation.

14.2 Tray towers

431

14.2 Tray towers Fig. 14.1A shows the important features of a typical cross-flow tray tower commonly employed for fractionation. Vapour and liquid are in countercurrent flow in the tower with vapour bubbles rising through the perforations on the horizontal trays to the tray above and liquid flowing downwards through the ‘downcomer’ connecting consecutive trays. However, on the individual trays shown, the two phases exhibit cross flow since the vapour emerging from the perforations rises through the liquid which flows across the tray from below the downcomer skirt at one end to the exit weir at the other end. It crosses the weir and enters the downcomer to the tray below. It is worth noting that the liquid flow area across a tray is nonuniform, increasing towards the central region. Mass transfer is achieved across vapour bubbles rising through the pool of liquid on the tray. Turbulence during contact promotes mass transfer and increases efficiency of contacting. However, beyond a limit of turbulence, entrainment of liquid droplets and froth carryover reduce stage efficiency. The vapour dispensers on the tray floor can have various layout patterns and various types of fitments as described in the following section. Besides vapour dispersers (perforations, bubble caps and valves), other components of the tray are as follows: -

TSR, a circular ring welded with the tower shell Bolting bars welded with tower shell Downcomer apron/skirt bolted on bolting bars Tray decks fitted on TSR and support beams Tray support trusses/beams which support tray segments

Tray towers are often retrofitted with intermediate packed sections mainly to increase capacity. The above description pertains to contacting trays, each of which provide one stage of vapoure liquid contact. In addition, the chimney trays, that are noncontacting trays, are used for side stream withdrawal from tray as well as packed towers as discussed later.

14.2.1 Contacting trays Contacting trays are classified based on arrangement for vapoureliquid contacting and liquid flow path on the trays. Countercurrent and cross-flow trays refer to liquid flow relative to the rise of vapour bubbles. Sieve tray without downcomer, also known dual-flow tray, is the only type of countercurrent tray. In this case, vapour and liquid pass through tray deck perforations in opposite direction and the tray occupies the entire tower cross section, thus providing slight vapour capacity advantages. There are no weirs, baffles or other attachments, and typically 15%e30% of the tower cross section is the total perforated area. The liquid forms a random pattern while draining and does not form a continuous streamlet from each hole. Liquid holdup and interfacial area are strong functions of vapour rate that limit the operating range (of vapour and liquid flow rates) with reasonable tray efficiency. These are also sensitive to load changes. Due to their simplicity and lower efficiency, this tray option although cheapest (on unit cross section basis) is used for small diameter columns and only when few stages are required. Additional trays are provided to take care of poor efficiency. A unique advantage of counterflow trays is their selfcleaning ability which is particularly advantageous in fouling, solid-contaminated and corrosive services.

432

Chapter 14 Column and column internals for gas

Cross-flow trays are the most common tray type and due to the longer liquid path has the highest tray efficiency. The simplest cross-flow arrangement, and consequently the cheapest to fabricate, is the single-pass tray Single-pass, multipass, cascade and reverse flow trays (Fig. 14.3A). It has a segmental downcomer at one side of the tray. The downcomer of consecutive trays is fitted on opposite sides such that liquid from the downcomer of the upper tray flows across the tray to the downcomer at the other end as shown in Fig. 14.3A. Since the flow across the tray is due to gravity, the depth of liquid reduces progressively along its travel path. A large gradient is undesirable as this leads to lower resistance to the flowing vapour at lower liquid depth (close to its downcomer). As a result, the vapour stream tends to flow preferentially or ‘channel’ from this zone instead of passing uniformly through the entire tray cross section. The gradient of liquid level/depth on the tray depends on liquid flow rate and length of liquid path. Towers that handle liquid with very low viscosity and low liquid to vapour flow rate have small liquid gradient per unit length of liquid path. Such towers use reverse flow trays (Fig. 14.3B), where despite the longer path, the gradient limit is not exceeded. In this case, the downcomers are all located on one side of the tower and the liquid is forced to flow around a central baffle at the far side which reverses its flow direction. The reverse flow tray provides more cap area at the expense of downcomer area and is advantageous only for very low liquid/vapour ratios. These trays are used in air separation towers. Requirement of reduction in gradient has led to development of multipass trays and cascade trays. These two types are encountered only in large diameter towers and are generally decided by the liquid loading. When liquid load is high with respect to vapour, multipass trays are suggested and for even higher liquid loads, the cascade type (with intermediate weirs) is recommended. Towers fitted with two-pass trays have different layout of consecutive trays. One has two downcomers at opposite sides and the other has a single downcomer located centrally as shown in Fig. 14.3C. This arrangement repeated in consecutive pair of trays splits the total liquid flowing on a tray into two streams which halves not only the liquid load per unit tray width but also the travel length of each liquid path. As a result, double pass trays provide considerably more liquid capacity and lower liquid gradient as compared to single-pass trays. However, they cost slightly more than a single-pass tray of the same diameter and due to the shorter liquid path, tray efficiency is lower as compared to single-pass ones. An odd number of passes is not recommended because of problems in liquid distribution. Higher number of passes, though possible, is not too common. Cascade trays have the tray deck in more than one level (usually two). The elevated deck and the lower deck have their individual exit weir to ‘step’ the tray floor at two elevations, thus further cutting down the liquid path in each section. Though the liquid flow rate on each deck is the same, reduction in travel path reduces the gradient. Fig. 14.3D shows a typical cascade tray. Tray types can be selected based on the information provided in Table 14.1.

14.2 Tray towers

(A)

(B)

(C)

Column shell

433

(D)

Side downcomer

Seal pan

Central downcomer

FIGURE 14.3 Cross-flow tray with the liquid flow path marked (A) single-pass, (B) reverse flow, (C) two-pass and (D) cascade tray.

Tray components: The different components of a tray are described in the following subsections.

Downcomer This is the conduit for the passage of liquid between trays. A level of liquid accumulated in the downcomer (downcomer backup) provides seal against vapour short circuiting through the downcomer passage. Typically, the design backup is about half the tray spacing. It is also important to ensure that the vapour bubbles disengage from the aerated liquid within the residence time in the downcomer and only clear liquid flows out to the lower tray. For liquids with foaming tendency, the downcomer backup should be kept low. Common downcomer types are as follows: • •

• •

Segmental downcomer e most common for medium to lager size towers (>300 mmf), as this provides maximum utilisation of tower area for downflow and results in high liquid flow capacity. Stepped apron downcomer e used to provide higher downcomer backup and increased active tray area. Narrower cross section at the lower portion provides extra resistance to liquid downflow resulting in higher downcomer backup. Pipe downcomer e usually for low liquid flows. It is cheaper than segmental downcomer in small diameter columns (2.5 m) towers

14.2 Tray towers

441

14.2.3 Tray construction Trays can be fixed or removable. Fixed trays are welded to the shell. These are difficult and expensive to maintain and are rarely used. Removable trays can be - Set of trays in the form of a cartridge which is lowered inside the tower from the top. As already mentioned, these do not perform very satisfactorily due to leakage through the shell clearance and are also expensive to maintain. - Trays fixed by flanges between tower sections. These are used for small diameter towers, usually D < 750 mm, and are expensive to fabricate and difficult to maintain. - For D > 1200 mm, trays are typically in sections that are taken inside through the manholes and assembled inside by bolting. Fitting of the valves/bubble caps is carried out inside the tower. Whenever sections are joined, bolts fixing the tray sections with the backing strip/truss have to be placed along the joining line. This reduces the number of valve/cap/holes that could otherwise be accommodated in that space. The shape as well as the size of the tray sections is carefully designed to accommodate maximum number of disperser elements. Tray manway section can typically be 450
500 mm or larger, depending on the manhole size on the tower. The second largest dimension of the tray section should obviously be such that it can be introduced through the manhole. Design value of hydraulic load is 600 N/m2 live load on tray þ 3000 N/m2 over downcomer seal area. An additional 1500 N concentrated load is imposed on any structural member during construction and maintenance.

14.2.4 Efficient operation of contacting tray Tower internals are designed to operate with maximum efficiency for specific vapour and liquid flow rates on individual trays. A properly designed tray must ensure the following: • • • •

Good vapoureliquid contact Sufficient liquid holdup for efficient mass transfer Sufficient area and spacing to keep entrainment and pressure drop within limits Sufficient downcomer area for unaerated liquid flow between trays

Net result of the closely coupled operation of all trays gives the performance of the tower. Maloperation of one tray affects the performance of the adjoining trays. Hence, the individual trays need to operate within ranges of vapour and liquid flow rate where tray efficiency as well as effect on the neighbouring trays is within acceptable (design) limits. The tray is designed for a specific pressure difference at the design point and is expected to operate with limited fluctuations around the mean value. Under this condition, the vapour flow is close to uniform through the openings on the tray. The range of satisfactory operation of a tray is usually represented on a two-dimensional plot with liquid and vapour flow rates as the abscissa and ordinate, respectively. This is referred to as ‘tray performance range diagram’. Fig. 14.8 shows the nature of diagram for typical sieve trays with the hatched area denoting the ‘area of satisfactory operation’. In order to arrive at the smallest tray size (diameter), the design point needs to be close to the upper limits of vapour as well as liquid flow rates (say at ‘A’ in the figure). This enables maximum utilisation of tray capacity and effectively results in a low tower diameter. Usually 10% to 15% overdesign with respect to vapour and liquid flow is considered as a reasonable design point.

E e n xce t r a ssi i n m ve en t

Chapter 14 Column and column internals for gas

G

Entra

inme

nt

B Satisfactory operation zone

A C

Weeping

Excessive weeping Dumping

D

F mer nco Dow oding flo

Vapour flow rate

442

E

Liquid flow rate

FIGURE 14.8 Typical nature of performance range diagram for a sieve tray with downcomer.

As shown in the figure, the range of satisfactory operation is bounded by the maloperations of excessive entrainment, downcomer flooding and entrainment flooding, excessive weeping and dump point. So in order to establish the range, a designer needs to understand the maloperations mentioned. Keeping the liquid flow rate (on the sieve tray) constant at ‘A’, if the vapour flow rate decreases below a limit, there is increased liquid flow through the openings. When the flow exceeds the minimum tolerable limit, the liquid level on the tray falls. The condition denoted as operating point ‘C’ is termed ‘weeping’. In this case, the entire liquid drains through the sieve holes before reaching the outlet weir. A further decrease of vapour flow leads to ‘excessive weeping’ say at operating point ‘D’ where the tray operation and contacting becomes unstable. At still lower vapour flow rates, there is nonuniform and fluctuating flow through the holes. The pressure drop across the tray fluctuates and is so low that it cannot support the liquid level. The liquid, under this condition, ‘dumps’ on the tray below and no liquid reaches the downcomer. The dumping condition corresponds to point ‘E’. Tray operation at conditions of excessive weeping and dumping results in poor quality of vapoureliquid contacting and a fall in tray efficiency below acceptable limit. At vapour flow rates above the design point ‘A’, the contact time between the phases decreases, consequently decreasing tray efficiency and increasing entrainment of fine liquid droplets to the upper tray. Intermixing of liquid carried from a lower tray to an upper tray reduces product enrichment, resulting in a lower separation efficiency of the tower. A small amount of liquid entrainment is unavoidable at normal throughput and the efficiency remains acceptable up to a limiting vapour flow rate. At higher vapour flows (say at point ‘B’), the vapour emerging from the perforations entrain substantial liquid droplets (and froth) and the efficiency drops sharply. This is the ‘entrainment flooding’ limit. The onset of flooding is detected by a steep increase in pressure drop and a sharp decline in efficiency. Keeping the vapour flow rate same, if the liquid flow rate is increased, the liquid level and the froth above it in the downcomer keep on increasing. The level on the tray also increases but to a lesser extent. At a limiting flow rate, the tray downcomer runs almost full of liquid. This is the condition of ‘downcomer backup flooding’. On further increase in liquid flow rate, the countercurrent flow of liquid

14.3 Tray design

443

entering downcomer and disengaging gas/vapour rising from the downcomer chokes the downcomer entry and liquid starts accumulating on the tray. This condition, described as ‘downcomer choke flooding’, is marked as ‘F’ in the figure. Exceeding this limit floods the entire tray and destabilises operation. This is characterised by froth carryover and pulsating flow of vapour through an undesirably high depth of liquid on the tray resulting in decreased vapoureliquid contact and reduced tray efficiency. At liquid flow rate lower than ‘A’, the liquid depth on the tray is less and liquid entrainment is more. The limiting range is marked as ‘excessive entrainment’ in the figure and corresponds to the design vapour flow rate with a much lower liquid flow rate. The onset of ‘excessive entrainment’ is marked as ‘G’ on the figure. Though the performance range diagram (Fig. 14.8) has been shown for sieve tray with downcomer, the limiting phenomena represented at the boundary of the shaded region are not very different for other cross-flow trays. However, the shape of the shaded area may differ as discussed in subsequent sections. The additional malfunctions for contacting trays not shown in Fig. 14.8 are as follows: (a) Pulsing occurs for (i) low and unsteady vapour rate, (ii) low slot opening ( QV;max , hso ¼ hs þ Height of shroud ring if the cap design has a clearance under the skirt and for caps flush with tray floor, the increased pressure drop is

 hso;overloaded Qv ;overloaded 2 ¼ . hso Qv ;max

Estimation of % slot opening is discussed in the following section. % Slot opening Slot opening at vapour load QV is obtained from Fig. 14.14 which presents slot opening (% of slot height in mm) as function of QV QV;max for different slot types (shape and size details in Table 14.6). The maximum vapour capacity QV;max in m3/s is calculated using Eq. 14.26. 1

QV;max ¼ 0:060478Cs As fhs ðrL  rV Þ=rL g2 2

3

(14.26) 3

for As, slot area per tray in m , QV;max in m /s, hs in mm and r in kg/m .

14.3 Tray design

461

100 Rs – trapezoidal slot shape factor, ratio of top to bottom width

Slot opening, % slot height

80

60 Rs

–

0.

00

(tr

Rs

ia

–

40

Rs

u ng

5 0.

–

la

r)

0

1.

00

(re

ct

a

u ng

la

r)

Maximum slot capacity formula: Qmax – CsAs

20

hsh

ρ – ρV L ρV

Rs

Cs

0.00 0.50 1.00

0.63 0.74 0.79

60

80

0 0

20

40

100

Vapor load, % maximum for fully loaded slots

FIGURE 14.14 Generalised correlation for slot opening.

pffiffiffiffiffiffiffi Cs in ðm=s mmÞ is 0.63 for triangular slot, 0.79 for rectangular slot and 0.74 for a trapezoidal slot with bottom width nearly twice the top width. hal , drop through aerated liquid, is expressed as hal ¼ bhds

(14.27)

b, the aeration factor, accounts for energy loss due to bubble formation, frictional resistance to flow through aerated mass, static head effects and difference between

 slot drop and slot opening, if any. It is 1=2

estimated from Eq. 14.28 from the value of FVa ¼ UVa rV

. UVa is the vapour velocity based on

3

active tray area in m/s and rV is the vapour density in kg/m . b ¼  0:0255
ðFVa Þ3 þ 0:1744
ðFVa Þ2  0:4282
FVa þ 0:9979

(14.28)

The dynamic slot seal, hds ¼ hss þ how þ D=2

(14.29)

hss is the static slot seal provided in Table 14.6, how , the height of liquid crest over weir is given by Eq. 14.14 and, D, the liquid gradient in mm is estimated from Eqs. 14.20e14.22. Typically, the mean dynamic slot submergence of 38 mm (1.500 ) is recommended for towers operating at atmospheric pressure. A mean submergence of 25 mm (100 ) is recommended for operation under vacuum, 50 mm (200 ) for 50e100 psig pressure and 75 mm (300 ) for 200e500 psig pressure conditions.

462

Chapter 14 Column and column internals for gas

Check for vapour distribution In bubble cap trays, a significant liquid gradient often causes vapour maldistribution resulting in low tray efficiency. This effect is checked from vapour distribution ratio rvd which is the ratio of liquid gradient D to the mean cap pressure drop ðhcap ¼ hdry þhso Þ. The normal design limit is recommended as ðrvd ¼ D = hcap Þ  0:5

(14.30)

In case this limit is exceeded, rvd can be minimised at the cost of increased pressure drop and reduced capacity by blanking off some of the caps or designing caps with higher pressure drop. Low rvd combined with low pressure drop and high capacity can be obtained by adopting wider cap spacing and higher skirt clearance to reduce D. Special care is required to calculate rvd for stepped caps and cascade trays. In both cases, the worst distribution on the tray is checked by calculating rvd for extreme pairs of rows. In addition, an effective D needs to be considered as the gradient between the corresponding rows minus the amount of stepping for stepped caps and the difference in dynamic submergence (hss þ how þ D to the particular row) for the rows in question for cascade trays.

Vapour velocity and corrected ‘approach to flooding’ Vapour velocity is given as the volumetric flow rate of vapour per unit tray net area or UV;n ¼ QV =An

(14.31)

Flooding % of the tray is recalculated with the ‘vapour velocity’ obtained from Eq. 14.31 and the tray flooding velocity (Eq. 14.7) expressed as sffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi ðrL  rV Þ UfV;n ¼ Csb rV The tray diameter needs to be increased in case flooding limit is reached, which does not happen often. The entrainment is also recalculated for the revised flooding % from Fig. 14.11.

Downcomer pressure drop (hdc;prdrop , mm of liquid) Downcomer head loss primarily occurs when the liquid flows out on the tray through the opening below the downcomer skirt. It is considered to be predominantly a function of velocity (m/s) of outflow below the skirt and is estimated using the following correlation   QL;m3 hr 2 hdc;prdrop ¼ 0:1275 (14.32) 100Ada where hdc;prdrop is in mm of liquid, QL;m3 hr is the liquid flow rate in m3/hr and Ada is the minimum flow area in m2, the lower value of area under downflow apron or between apron and inlet weir (if provided).

14.3 Tray design

463

Downcomer backup (hL;dc , mm of liquid, for all cross-flow trays) Downcomer backup in terms of clear liquid level is expressed as hL;dc ¼ hw þ how þ hdc;prdrop þ htray þ D

(14.33)

All terms on the right-hand side of Eq. 14.33 are discussed above. An important check is hL;dc  TS=2

(14.34)

Velocity and residence time in downcomer The two considerations for sizing of downcomer are (i) residence time and (ii) maximum allowable velocity. Residence time in downcomer tdc is calculated from liquid flow rate and clear liquid volume in downcomer as tdc ¼

Adc hL;dc QL

(14.35)

The minimum residence time needs to lie within 3e7 s, the higher limit is for foaming systems. The maximum velocity of clear liquid in the downcomer is limited by choking in the downcomer and disengagement of vapour bubbles from the liquid. The recommended range of maximum downcomer velocity is 0.06e1.5 m/s (0.2e0.5 ft/s) with the typical value being around 0.1 m/s.

Downcomer throw over the weir Liquid throw over weir (dtw ) is the distance from the top of the weir to the level of aerated mass in downcomer. Throw of liquid over weir into downcomer is not important in case of segmental downcomers. For multipass trays, it is important for centre and off centre downcomers where the two liquid streams from opposite weirs are thrown towards each other. However, this is not important for the side downcomers. In case of central downcomer, dtw is estimated as pffiffiffiffiffiffiffiffiffiffiffiffiffi dtw ¼ 0:8 how hff (14.36) where hff the height of free fall in downcomer measured from weir is hff ¼ TS þ hw  hfd

(14.37)

hfd ¼ hL;dc =rrel

(14.38)

rrel ¼ 0:5

(14.39)

and If dtw calculated from Eq. 14.36 is greater than half the width of centre downcomer, antijump baffles are employed.

464

Chapter 14 Column and column internals for gas

System (foaming) factors (applicable for all cross-flow trays) The flooding of a tray due to excessive downcomer backup of liquid along with foam or due to entrainment of foam with vapour leaving the tray are often different for specific systems and applications. In either case, this is reflected as loss of tray efficiency that needs to be taken care of while using the generalised correlations presented so far. Traditional flooding equations are corrected with a (derating) system factor (SF). Actual downcomer backup (including foam) is calculated by dividing the calculated backup (hL;dc ) by (SF). Some designers also discount flooding velocity (UfV ), calculated using SouderseBrown type of equations with the system factor (SF). System factors of some common services are listed in Table 14.9 and may be used for the pertinent cases. Table 14.9 System factor for different services. SF

System Crude/vacuum column

0.85e0.9

Crude pretopping column

0.8e0.85

H2S/CO2 (gas)- amine absorber

0.65e0.7

Glycol absorber

0.65e0.7

Amine regeneration column

0.85

Glycol regeneration column

0.85

Sour water stripper column

0.7

Caustic regeneration

0.3e0.6

Oil absorbers

0.85

Weep holes These have a diameter of 3e5 mm. A general recommendation is to provide 275e280 mm2 of weep hole area per m2 (4 square inch per 100 ft2) of net open liquid tray area summed over all trays in the tower. Alternatively, the following empirical expression can be used for bubble cap tray draining:  ð10:895N þ 1:36192Þ m0:12 An;L l q¼ (14.40)  1:2 deq;wh =h0 r0:25 l where N is total number of actual trays in column, q is the draining time of the tower in minutes. Average viscosity and density of liquid at temperature in tower is ml cP and rl gm/cm3. deq;wh (mm) is the diameter of a circle whose area is same as the total area of all weep holes on one tray. h0 (mm) is the lower of the bubble cap riser height and the height of overflow weir. An;L (m2) is the net open liquid area of one tray. The accuracy of the relationship is 6% on an average. During design, the draining time is set to a reasonable value e say typically 6e8 hours and deq;wh is calculated using Eq. 14.40. Number of weep holes is then calculated based on a chosen weep hole diameter.

14.3 Tray design

465

14.3.2 Sieve tray design (cross-flow type e with downcomer) A sieve tray comprises of a flat metal sheet perforated with round holes and suitably supported in the tower. The tray is connected with one or more downcomers for liquid discharge and may contain weirs and baffles for directing vapour and liquid flows. Fig. 14.15 depicts the schematic of the sieve tray. A comparison with Fig. 14.9 denotes similar liquidevapour contacting action of a sieve tray and a bubble cap tray, both having sufficient liquid submergence to prevent short circuiting of vapour. There are, however, two differences between vapour flow in sieve tray and bubble cap trays. In sieve trays, the vapour emerging from the sieves flow primarily in the vertical direction unlike the tortuous path followed during flow through bubble caps and since there is no built-in liquid seal, only vapour flow can prevent liquid weeping through the holes. The active tray area (Ao,) is characterised by the aerated mass or froth of height hf equivalent to an effective hydrostatic head hL . The aerated mass collapses in the calming section upstream of outlet weir. The calming section may be provided for partial froth collapse but is generally not needed and the overflow material is in reality aerated. The equivalent height of clear liquid hlo at the liquid exit end is calculated as the sum of the outlet weir height and the crest of equivalent liquid falling over the weir. Downcomer design must take into account secondary froth formation and allow space for froth collapse; otherwise froth density may be too low for adequate liquid outflow resulting in decrease of efficiency and increase in pressure drop. The terms ‘foam’ and ‘froth’ are often interchangeably used in literature. In this book, froth is the expanded material formed during passage of gas or vapour through a liquid and if the expansion is related more to liquid physical properties than to method and degree of aeration, the material is ‘foam’. In sieve tray design, ‘foamability’ is important and taken into account. Design deliverables for cross-flow sieve trays are as follows: • • • • • • • • • •

Liquid flow arrangement Active area Area occupied by perforations Hole size, pitch and arrangement Hole blanking, if required Tray baffles and calming zones Downcomer area, type and clearance Tray inlet arrangement Outlet weir e type and dimensions Tray thickness, material and tray levelness

Steps of design

    Input: mL ðkg =sÞ; mV ðkg =sÞ; rL kg m3 ; rV kg m3 ; j%; sðdynes=cmÞ; TSðmmÞ Initial guess: d0 ¼ 5 mmð3 =1600 Þ; A0 =Aa ¼ 0:1; TS ¼ 450=600 mm; Tray thickness: 2 mm; straight weir, hw ¼ 50 mm, lw =D ¼ 0.77; segmental downcomer with Adc/A w0.12

466

Chapter 14 Column and column internals for gas

End wastage

Ao, tray

Adc

Adc

Calming zone

Wdc

Dφ

lw

Foam level Δ

Perforations TS

hw

hdc, clearance

hL, dc

how

FIGURE 14.15 Schematic representation of sieve tray dynamics.

(1) Tray type and spacing Single-pass cross-flow tray is selected from economic considerations even for large diameter columns. TS of 300e410 mm (1200 e1600 ) is common. Higher tray spacing up to 760 mm (3000 ) may be adopted for high vacuum services and it is rarely less than 230 mm (900 ). Usually TS for sieve plates can be 150 mm (600 ) less than for a corresponding bubble cap tray. (2) Estimation of tower diameter Sieve tray diameter is estimated from ‘incipient flooding’ limit which corresponds to maximum tray capacity. Although both liquid or vapour capacity limitations can lead to ‘incipient flooding condition’ and both are limiting at true flooding, the former is more common and should be checked initially. The procedure presented in Section 14.3.1 can be followed and Fig. 14.16 can be

14.3 Tray design

467

used to estimate K (analogous to Csb in Eq. 14.7) as function of liquidevapour flow  the constant rffiffiffiffiffiffi 1 m L rV parameter: FLV ¼ . mV rL rffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi rV (14.41) K1 ¼ Uf V;n ðrL  rV Þ Fig. 14.16 is based on small diameter perforations (do < 6:5 mm). Larger perforation size results in higher entrainment. The correlation is valid under the following conditions: (i) ðAo =Aa Þ  0:1 and (Adc/A) ¼ 0.12. (ii) ðhw =TSÞ < 0:15 (iii) s ¼ 20 dynes=cm (iv) Nonfoaming systems or systems with low foaming tendency In case of foaming systems, the vapour velocity Uf V should be 75% of the value predicted following the aforementioned procedure. For liquid surface tension ss20 dynes=cm and ðAo =Aa Þ  0:1, the corrected value of K1 is ¼ K1 ðs=20Þ0:2 KCorrected 1

(14.42)

If ðAo =Aa Þh0:1 and ss20 dynes=cm, the corrected value of K1 is ¼ ½5ðAo = Aa Þ þ 0:5K1 ðs=20Þ0:2 KCorrected 1

(14.43)

100

Plate spacing.m –1 K1 10

0.90 0.60 0.45 0.30 0.25 0.15

10–2 0.01

0.1

FLV

FIGURE 14.16 Flooding capacity e sieve trays.

1.0

5.0

468

Chapter 14 Column and column internals for gas

The corresponding vapour velocity at flooding, Uf V;n , based on An is given by sffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi ðrL  rV Þ Uf V;n ¼ KCorrected 1 rV

(14.44)

Similar to bubble cap trays, An, is the net area of tray available for liquid disengagement and is typically (A e Adc). Unusual baffling can reduce this area. If a splash baffle is used at the outlet weir, An ¼ Aa and the velocity obtained from Eq. 14.44 is Uf V;a . Usually 80%e85% approach to flooding is used. The calculated tower diameter D is rounded off based on Table 17.8 as discussed in Section 14.3.1. The results of flooding analysis give a tentative diameter-tray spacing combination which may be adjusted in subsequent design calculations. For significant variations in volumetric flow rates of liquid and vapour, the limiting tower size is obtained by making separate flooding calculations for different points in the tower. Presence of side stream draw and circulating refluxes may warrant such situations in distillation columns. It is usually sufficient to make two estimates e above and below the feed point and changes in vapour flow rate are adjusted by blanking off some rows of holes, while liquid flow rate variation is addressed by adjusting liquid downcomer area. Different column diameters for different tower sections are adopted only for significant variations in flow rates and are economical only for large towers. This is in fact rare for sieve tray towers. (3) Tray layout Sieve trays with perforation diameter do z3  12 mm (1/800 e1/200 ) are commonly employed. The recommended size for nonfouling services is 5 mm (3/1600 ) and for fouling liquids and liquids containing solids, larger perforations (12.5 mm/0.500 ) are adopted. For vacuum services and systems with low Perforation size. surface tension, do z3  6 mm (1/8e1/400 ). Air separation towers are typically equipped with sieve trays of 1 mm or less perforation size. Perforations are either drilled or punched on the trays. Punching is done for 2:5  do  19 mm. The direction of punching oriented in the direction of vapour flow is often preferred. The minimum perforation size is generally limited by dimension of punch die which depends on thickness and material of tray (a general thumb rule e perforation size should not be less than sheet thickness for C-steel and 1.25 times the sheet thickness in case of stainless steel). In usual practice, tray thickness of 12 or 14 US standard gauge is employed except for C-steel where 10 gauge thickness is used. Thus, typical perforation size is 3.2 mm (1/800 ) in 14 US standard gauge tray thickness of 2 mm (0.07800 ) and 4.76 mm (3/1600 ) in 10 US standard gauge SS tray of thickness 3.6 mm (0.14100 ). Interestingly, at small perforation diameter (do  2 to 3 mm), uniform gas flow through all holes is the limiting mechanism and the minimum gas load decreases with increasing perforation size while for do > 2 to 3 mm, liquid weeping through tray becomes decisive and the minimum gas load increases with increase in perforation diameter.

14.3 Tray design

469

As a preliminary estimate, pitch P is calculated by considering the net perforated area discussed above. If the calculated pitch does not correspond to a commercial standard, the next smaller pitch is selected and the (net) perforated area is recalculated to determine the amount of blanking (perforations covered with sheet metal) required. Blanking strips Hole pitch and number should be distributed uniformly over active area except when used for calming zones at the weirs. Blanked area should normally not exceed 25% of active area (to avoid excessive entrainment). In order to avoid ‘dead spots’ in the active area, the width of the backing strip should not exceed 70% of tower diameter for small towers and 50% for larger ones The active tray area usually encircles the perforated area at a distance of 50e75 mm (2e300 ) from the peripheral perforations and the perforations are arranged on a triangular pitch with liquid flow normal to the rows. Generally, 2:5d0  P  5d0 with P ¼ 3:8do commonly preferred. For effective tray action, perforations should not be more than 65e130 mm (2.5e300 ) apart. 6 pffiffi As mentioned in Section 14.3.1, the number of holes per m2 of net perforated area is 2
10 for 3 P2 triangular pitch, when P is in mm. The net perforated area Ao;tray is the difference between active area Aa and area covered by (i) tray supports, tray rings, etc., (ii) inlet and exit calming zones and (iii) end wastage area. Generally, the entry and exit calming zones are 75 mm each for D < 1:5 m and 100 mm for D > 1:5 m. Perforation area The acceptable range of ðAo =Aa Þ is 0.05e0.15 with the preferred value being 0.1 and for pressure services, the range is 0.08e0.10. The ratio is never less than 0.05 and for some critical services, the ratio can be as high as 0.17 to 0.175, provided weeping does not occur. Open area ratio Ro is used only to identify perforated sheet metal material and for a given section of material, it refers to the ratio of hole area to total area. For equilateral triangular pitch, Ro ¼

A0 ¼ 0:905
ðdo =PÞ2 A

(14.45)

and for square pitch Ro ¼ 0:785ðdo =PÞ2

(14.46)

Perforated sheet metals usually come in standard width of 900 mm (3600 ) and 1200 mm (4800 ) and in standard lengths of 2500 mm (9600 ), 3000 mm (12000 ) and 3500 mm (14400 ). For special requirements, other widths and lengths can be provided. Tray support members can be spaced at 300 mm (1200 ), 400 mm (1600 ), Tray supports 450 mm (1800 ) or 600 mm (2400 ). Using blank margins of perforated sheets as tray support minimises wastage of metal. Generally, tray spacing and support members employed with sectional trays need to consider accessibility through manways. The width of support ring is 50e75 mm and the support ring does not extend into the downcomer area. (4) Area ratios and adjustment of flow conditions The tower cross-sectional area A ¼ ðp=4ÞD2 is reestimated from the rounded off D for single-pass cross-flow tray area and the area of the individual tray components are calculated as Adc ¼ 0:12A, An ¼ 0:88A Aa ¼ 0:76A; Ao ¼ 0:1Aa . Based on these, the flow conditions are adjusted and UV; n is calculated from Eq. 14.9 to check that j is within permissible limits (80%e85%).

470

Chapter 14 Column and column internals for gas

(5) Entrainment Fractional entrainment j (kg moles/kg moles gross liquid flow) is determined from Fig. 14.17 corresponding to FLV and % flooding and the total amount entrained is obtained from Eq. 14.11. j in the range 0.1e0.2 represents optimum conditions.

100 9 8 7 6 5 4 3

Percent flood 2

95 10–1 9

90

Fractional entrainmentψ

8 7 6

80

5 4 3

70

2

60 50

10–2 9

45

8 7 6

40

5

35

4 3

30 2

10–3

2

3

4

5 6 78 9

10–2

2

3

4

5 6 7 8 9

10–1

100

FLV

FIGURE 14.17

qffiffiffiffi  rV mL Effect of flow parameter FLV ¼ m on fractional entrainment j for sieve tray. r V L

From Smith, B.D. (1963). Design of equilibrium stage processes. McGraw-Hill.

14.3 Tray design

471

Fig. 14.17 shows that at high mL =mV , sieve trays may be operated quite close to flooding point before significant entrainment occurs. Accordingly, optimum values lose their significance under such conditions since flood point itself cannot be predicted accurately. (6) Tray pressure drop The pressure drop (htray ) across the tray in mm of liquid is htray ¼ ho þ bðhw þ how Þ

(14.47)


103 h

and in pressure units (Pa), Dptray ¼ 9:81 tray rl . ðhw þ how Þ is the operating liquid seal at tray outlet weir in terms of clear liquid height. hw is typically 50 mm. Table 14.19 can be referred for selection of hw under different pressure conditions and how can be estimated from Eq. 14.14. ho , the head loss due to vapour flow through perforations (mm of liquid) is ho ¼ 51
ðrV = rL Þ
ðUVo =Co Þ2

(14.48)

UVo , the vapour velocity (m/s) through the perforations corresponding to ho , is calculated as UVo ¼

ðmV =rV Þ Ao

And the orifice coefficient Co is evaluated as

m values for different

Tray thickness Hole diameter





Ao Co ¼ 0:7205
Aa

(14.49)

 þm

(14.50)

are listed in Table 14.10.

Table 14.10 Evaluation of orifice coefficient Co . 

Tray thickness Hole diameter



m

1.2

0.8142

1.0

0.7736

0.8

0.7080

0.6

0.6733

0.2

0.6404

 0:1

0.5885

472

Chapter 14 Column and column internals for gas

1=2

The aeration factor b in Eq. 14.47 is obtained from Eq. 14.51 as a function of FVa ¼ UVa rV as b ¼ 0:5792 þ 0:4027
eð1:5806
FVa Þ

(14.51)

where UVa is the vapour velocity based on active tray area in m/s and rV is the vapour density in kg/m3. The desirable range of the variable in the above equation is 0:305  Fva  3:05. For j > 0:1, two-phase flow increases pressure drop through perforations and the equivalent head loss hj>0:1 is given by 
  j FLV hj>0:1 ¼ hdry 1 þ 15
(14.52) 1j Alternatively, the hydrostatic head of aerated mass on tray, hf , corresponding to clear liquid height hL (in mm) can also be estimated from relative froth density from the expression hf ¼ hL =rrel

(14.53)

rrel ¼ hL =hf ¼ 2b  1

(14.54)

where the relative froth density

Typically, rrel z0:4  0:7 and for design purposes, rrel ¼ 0:5 is normally assumed. The same value of rrel is adopted for downcomer which needs to respect hL;dc  f0:5
ðTS þ hw Þg

(14.55)

(7) Weeping Vapour velocity at weep point is the minimum velocity for stable operation and the perforated area is chosen to ensure that the vapour velocity is above weep point at the lowest operating flow rate. Weeping occurs on the trays due to the (static) head of (equivalent clear) liquid hL and is opposed by liquid surface tension and vapour flow through perforations. This gives the condition to prevent weeping as ðhs þ ho Þ > ðhlo ¼ hw þ how Þ

(14.56)

where the maximum depth of liquid hs (in mm) that can be sustained by surface tension is empirically given by hs ¼

15:806s rL do

(14.57)

for rL in kg/m3, s in dynes/cm and do in mm. Tray weeping may occur at part or whole of the tray. This requires the parameter ðAo =Aa Þ to be considered along with hlo and the conditions to prevent weeping are given by the following criteria for 0.06 < ðAo =Aa Þ < 0.14 and ðAo =Aa Þ ¼ 0:2. (a) When 0:06 < ðAo =Aa Þ < 0:14: n o (14.58) ðho þ hs Þ >  7.433
104
ðhlo Þ2 þ 0.2358
hlo þ 3.52

14.3 Tray design

(b) When ðAo =Aa Þ ¼ 0:20: n o ðho þ hs Þ >  7.655
104
ðhlo Þ2 þ 0.43676
hlo þ 5.3

473

(14.59)

Since the influence of weeping on tray efficiency depends on the fraction of total liquid downflow that weeps, even a small amount of weeping can be relatively serious at low liquid flow rates. Alternately, the minimum design vapour velocity to prevent weeping can be obtained as UV0;min ¼

K2  0:90ð25:4  d0 Þ r0:5 V

(14.60)

where UV0;min is the minimum vapour velocity through the holes (based on hole area) in m/s, d0 is hole diameter in mm and K2 is a constant, expressed as a function of clear liquid depth on tray ðhlo ¼ hw þ how Þ as K2 ¼

32:9ðhlo Þ0:62 1:044 þ ðhlo Þ0:62

(14.61)

Eq. 14.61 is valid for 14  hlo  111:5 (8) Liquid gradient across tray ðDÞ Liquid gradient problems are much less severe in sieve tray as compared to bubble cap and valve tray design since the resistance to liquid flow is smaller in this case. So large diameter towers without double flow paths are possible. In normal practice, the effect of gradient is neglected unless D > 19 mm (0.7500 ). Nevertheless, for long flow paths and high liquid rates, liquid gradient on sieve trays needs to be checked. It is also significant in vacuum operations where low weir height causes D to be a significant fraction of total liquid depth. Similar to the case of bubble cap trays, the condition for stable tray operation is given by Eq. 14.30 as ðD =ho Þ < 0:5 where ho is given by Eq. 14.48 and D in mm of clear liquid is calculated from Eq. 14.62 with all units in m. D¼

f
lpath
Uf2 9:81
Rh


103

lpath ¼ D  wdc;inlet  wdc;outlet

(14.62) (14.63)

where wdc;inlet and wdc;outlet are the width (weir to wall) of the inlet and outlet downcomer and D is the tray diameter in single-pass trays. In case of double pass, the appropriate geometry is considered as discussed in Section 14.3.3. Rh , hydraulic radius for cross flow of aerated mass is given as Rh ¼ hlo
lf =ð2hlo þ lf Þ

(14.64)

The average liquid flow path width for a single-pass tray is the average of tray diameter D and weir length lw , lf ¼ ðD þ lw Þ=2

(14.65)

474

Chapter 14 Column and column internals for gas

  Rh U f r l The friction factor f is obtained from Reynolds number Re ¼ m for a known weir height l hw as f ¼ expð  1:0583
lnðReÞ þ cÞ

(14.66)

f is estimated by interpolating values obtained from Eq. 14.66 for known Re at two nearest hw and the corresponding c is available in Table 14.11. Table 14.11 Evaluation of friction factor f for known Re and hw . hw

c

10.2 mm (0.4)

6.5411

17.8 mm (0.70 )

6.9849

25.4 mm

(10 )

7.4102

38.1 mm

(1.50 )

7.8915

50.8 mm

(20 )

8.2246

76.2 mm

(30 )

8.7058

97.6 mm (40 )

9.1054

Using consistent units, the average liquid velocity on the tray in m/s is estimated as Uf ¼ QL =ðhlo
lf Þ

(14.67)

The calculated D is checked for the acceptability condition mentioned in Eq. 14.30. For significant D, perforations nearer to tray inlet weep before those closer to tray outlet. (9) Downcomer dynamics Both sieve tray and bubble cap tray employ the same downcomer design and the only difference in liquid handling occurs in the aerated zone. Downcomer backup in terms of clear liquid is calculated by adding the head loss terms as discussed in Section 14.3.1. hL;dc ¼ htray þ hw þ how þ D þ hdc;prdrop

(14.33)

Evaluation of the terms on the right-hand side of Eq. 14.33 is discussed above. hL;dc is measured from plate surface, how and hdc;prdrop are estimated from Eqs. 14.14 and 14.32 provided in Section 14.3.1 (bubble cap tray design). An important check is hL;dc  TS=2

(14.34)

Liquid throw over weir hfd is important only for multipass trays which is not common in sieve trays. Antijump baffles can be installed to ensure that liquid flows smoothly into the central downcomer.

14.3 Tray design

475

14.3.3 Valve tray design Valve trays (also known as ballast trays from M/s Koch-Glitsch) are essentially sieve plates with larger diameter of perforations covered by flaps which get lifted with increasing vapour flow. These are mostly proprietary items. The procedure presented here is by M/s Koch-Glitsch that has been primarily from the Ballast Tray Design Manual (Bulletin 4900), sixth edition. The design procedure estimates tray diameter and spacing, capacity and pressure drop. Subsequently, changes in diameter, tray spacing, cap spacing or downcomer specifications are made to arrive at an optimised design with maximum capacity, maximum efficiency and minimum cost. The procedure outlined claims neither a too conservative nor a too tight design. The correlations used in the design procedure outlined below has been developed by Carl Branan (The Process Engineer’s Pocket Handbook, Vol. 1, Gulf Pub. 1990) Design steps: 1. Approach to flooding (j %) is taken as not more than 82% for normal services and not more than 77% for vacuum operations. Typically, 65  j  75 for D < 0:9 m. A thumb rule is j ¼ 70% for D  2 m and j ¼ 80% for D > 2 m 2. Tray spacing, TS, typically is 600 mm (2400 ) 3. Estimation of approximate tower diameter (D) for calculation of flow path length The approximate tower diameter Dapprox in m can be estimated from the following set of equations at known actual vapour and liquid load of Qv and QL m3/s respectively. ðDapprox Þ2 ¼ 9:925
A þ 624:03
A
B þ 630:31
A
B2 þ 433:11
ðBÞ1:719 Where for single pass tray, B ¼ QL in m3 =s and A ¼ Qadj v ¼ Qv

rffiffiffiffiffiffiffiffiffiffiffiffiffiffi rv rl  rv

(14.68)

(14.69)

Eq. 14.68 is valid for single pass trays for 600 mm (2400 ) TSat 80% flooding condition. Estimation accuracy with respect to the original nomogram is 15% for 600  D  1200 mm and for higher diameters up to 3000 mm the accuracy is w5%. The range of validity of flow rates are: Qv < 0:85 m3/s and QL < 0:095 m3/s. The same equation can be used for multipass trays using the values of A and B (in m3/s) from Table 14.12. Table 14.12 A, B to be used in Eq. 14.68 for estimation of tray diameter Number of tray passes

Single

2 pass

4 pass

Qadj v

Qadj v =2

Qadj v =4

B(m /s)

QL

QL =4

D(m)

Dapprox from Eq. 14.68

QL =2 pffiffiffi 2
Dapprox obtained from Eq. 14.68

A(m3/s) 3

2
Dapprox obtained from Eq. 14.68

476

Chapter 14 Column and column internals for gas

4. No. of tray passes as function of tower diameter is decided based on Table 14.13.

Table 14.13 Tray passes in valve tray towers. Number of tray passes, NP

Dmin in m (ft)

Preferred D in m (ft)

2

1.52 (50 )

1.8 (60 )

3

2.44

(80 )

4

3.05 (100 )

3.66 (120 )

5

4 (130 )

4.5 (150 )

2.74 (90 )

5. The approximate diameter and number of tray passes NP is used to calculate flow path length, FPL in mm from the following expression involving Dapprox in m. FPL ¼

750
Dapprox NP

(14.70)

6. Minimum active area of tray (Aa;min m2 ) Aa;min ¼

3:2808
Qadj v þ ðQL
FPLÞ=224:24 CAF
ðj=100Þ

(14.71)

3 for Qadj v and QL in m /s and FPL in mm, j is % approach to flooding specified in Step 1 of design, FPL in mm is calculated from Eq. 14.70 and vapour capacity factor (CAF) is estimated as function of tray spacing TS (in mm) and vapour density rV (in kg/m3) from Eq. 14.72. For rV < 2.7 kg/m3 (rV < 0.17 lb/ft3),

1=6

CAF ¼ 6:412
103
SF
TS0:65
rV

(14.72a)

and for rV > 2.7 kg/m3 (0.17 lb/ft3)   CAF ¼ SF
0:1392 þ ðTS = 1305:2Þ  TS2 = 2560:2
103  ðTS
rV Þ = 640.74
103 (14.72b) System factor SF in Eq. 14.72 can be obtained from Table 14.14.

14.3 Tray design

477

Table 14.14 System factor for different services. Tray SF

Service

Downcomer SF

Nonfoaming, regular systems

1.00

1.00

Fluorine systems, e.g., BF3, Freon

0.90

0.90

Moderately foaming, e.g., oil absorbers, glycol and amine regenerators

0.85

0.85

Highly foaming, e.g., glycol and amine absorbers

0.73

0.73

Severely foaming, e.g., MEK units

0.60

0.60

Foam-stable systems, e.g., caustic regenerators

0.30 to 0.60

0.3

CAF increases with increase of TS and decrease of rV . However, for a particular rV , there is a limiting TS beyond which CAF does not increase with tray spacing. This limiting TS in mm is  given by Eq. 14.73 for 32  rV  86:5 kg m3 TSlimit ¼  63:284
104
r3V þ 1.4861
r2V  121.62
rV þ 3800:2 (14.73)  3 In cases of rV < 32 kg m , TSlimit ¼ 1200 mm. 7. Downcomer design velocity Udc;design (in m/s) is the maximum velocity of clear liquid in the downcomer. It is limited by downcomer choking and disengagement of vapour bubbles from liquid. It is taken as the smallest Udc;design value obtained from the following three equations: Udc;design ¼ 0:17
SF pffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi Udc;design ¼ 6:957
103
ðrL  rV Þ
SF pffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi Udc;design ¼ 2:525
103 TS
ðrL  rV Þ
SF

(14.74a) (14.74b) (14.74c)

where rL & rV are liquid and vapour density in kg/m3, TS is tray spacing in mm and SF is specified as downcomer SF in Table 14.14 for different services. Typically, Udc;design is recommended as 0.061e0.15 m/s (0.2e0.5 ft/s) and minimum downcomer residence time is 3e7 s. 8. Minimum downcomer area (Adc;min m2 ) is given by  (14.75) Adc;min ¼ QL = Udc;design
j = 100 is in m/s and j is the approach to downcomer flooding in % where QL is in m3/s and Udc;design If Adc < 0:11
p
D2 4 , i.e., below 11% of tray area, then Adc;min is the smaller of (i) 2
Adc  calculated from Eq. 14.75, or (ii) 0:11
Aa;min , Aa;min estimated in Step 6. 9. Minimum tower cross-sectional area (Amin m2 ) is the larger one of that given by Eq. 14.76a and b. Amin ¼ Aa;min þ 2
Adc;min

(14.76a)

478

Chapter 14 Column and column internals for gas

and Amin ¼ 0.0929


Qadj v 0:78
CAF
ðj=100Þ

! ¼

0:1191
Qadj v CAF
ð j=100Þ

! (14.76b)

10. Minimum tower diameter D in m is recalculated from Amin obtained from Eq. 14.76 as pffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi D ¼ 4
Amin =p (14.77) D is rounded off as per Table 17.8 and the number of passes are finalised accordingly from Table 14.13. The minimum cost tower will have cross section of Amin . 11. Downcomer width and area Layout of tray areas with 1e4 passes with the downcomer width (H) and FPL marked are shown in Fig. 14.18. The downcomer area at the top is divided in proportion to the liquid rate received and the active area served. In case of multipass trays, the liquid rate to the weirs in m3/hr per m of weir length is set as equally as possible. Also the downcomer lengths are slightly adjusted to modular flow path lengths. From Fig. 14.18, FPL can be expressed as FPL ¼

D  ð2H1 þ H2 þ 2H3 Þ NP

(14.78)

FPL cannot be less than 405 mm (1600 ) in case of trays with manways. Downcomer area (Adc, m2) is given as Adc;i ¼ Hi
SF
D

(14.79)

where SF is the span factor fraction defined as the wall to wall distance at the midpoint of downcomer expressed as fraction of tower diameter. Approximate values of SF for two, three and four pass trays are shown in Tables 14.15 for H2 and H3 marked in Fig. 14.18. Table 14.15 Span factor (SF) for different downcomers SF H2

# of passes

H3

2

1

-

3

-

0.95

4

1

0.885

Refer to Fig. 14.18 for nomenclature; H1 from weir length.

14.3 Tray design

H1

H1

FPL

H1

Single pass

H1

FPL

H3

FPL

H2

479

H1

FPL

Two pass

FPL

H3

FPL

H1

H1

Three pass

FPL

H3 FPL H2 FPL

H3

FPL

H1

Four pass

FIGURE 14.18 Layout of tray passes.

12. Active area (Aa ) Aa (m2) is the area available for ballast units between the inlet and outlet edges of the tray for straight or sloped downcomers. Aa ¼ A  ð2Adc1 þ Adc3 þ 2Adc5  2Adc7 Þ or Aa ¼ A  2Adc;avg

(14.80)

where Adc;avg is the average Adc for odd and even trays. Sloped downcomers are used with recessed inlet areas or draw sumps.

480

Chapter 14 Column and column internals for gas

13. Number of valves on a tray Actual number of valves on the tray active area can only be found from a detailed layout of the tray that considers the location of the major beam and truss lines. Truss lines are usually parallel to the liquid flow direction but in case of trays with major beam, the truss lines are perpendicular to the liquid path. The approximate number of valves accommodated on a tray is thus the number of rows multiplied by the average number of valves per row, with correction for the tray manway loss. This is found as (a) When truss lines are parallel to liquid flow path (for towers not having a major beam)   FPL  0:216 þ 1
NP (14.81a) No of rows ¼ 0:5
Base where FPL and Base are expressed in m Valves per row ¼

WFP  0:8
ðNo. of major beams þ 1Þ 0:146
NP

(b) When truss lines are perpendicular to liquid flow path (with a major beam)   FPL  0:0445
ðNo: of TrussesÞ  0:15 No of Rows ¼
NP 0:0635 Valves per row ¼

1000
WFP  2
ðNo. of major beams þ 1Þ Base
NP

(14.81b)

(14.81c) (14.81d)

where FPL is the flow path length in m WFP is the width of flow path (in m) ¼ ðAa Þ=FPL for Aa in m2 NP is the no of passes, and Base ¼ base spacing of valves (in mm) which is the centre-to-centre distance between valves in a row. Usually Base ¼ 76 mm (300 ). It can also be 88.9 mm (3.500 ), 101.6 mm (400 ), 114.3 mm ( 4.500 ), 152.4 mm (600 ). Spacing 00 For traditional round valves, orifices of 39 mm (1 17=32 ) diameter are punched in the deck in a triangular pattern. The base pitch of the triangle (parallel to liquid flow direction) varies from 76 00 to 152 mm (300 to 600 ). The triangle height is typically 63.5 mm (2 1=2 ) except across truss lines (joints). Truss lines are preferably parallel to liquid flow, but may be perpendicular to liquid flow in larger towers that have at least one major beam. The valve base pitch is inversely proportional to the required valve quantity, which is a function of the design vapour rate. There will be approximately 130e150 valves per square meter of active area when using a base pitch of 76 mm (300 ). A typical layout for a 1.5 m (50 ) single pass tray is shown in Fig. 14.19.

481

1.25

4.25

2.5

11.75

9.69

14.3 Tray design

17.06

MANWAY

9.69

11.75

FLOW

25 × 1.50 = 37.50

7.00

11.25

4.25 TWR I.D. 5 FT- 0.00 IN 42.69

8.69 1.25

8.69 0.81

FIGURE 14.19 Typical layout for 50 single-pass tray. Courtesy of Koch-Glitsch LP, Wichita, Kansas.

482

Chapter 14 Column and column internals for gas

14. Weir height is usually 50 mm (200 ) in most services. Generally, a weir height less than 19 mm (0.7500 ) is not recommended but in vacuum towers, it can be as low as 12 mm (0.500 ). For a high residence time, e.g., a system involving chemical reaction, hw can be 150 mm (600 ). For hw > 0:15TS, the effective tray spacing to calculate% flooding is the TS reduced by the excess weir height over 0:15TS. 15. Pressure drop (hdry mm of liquid) Dry tray pressure drop is the pressure drop of vapour in passing through the valves on the tray in dry condition, similar to that of vapour flow through an orifice, where the head loss is proportional to the square of velocity (UV;valve m/s) through the orifice. It is taken to be the larger of the following two values calculated: (a) Units part open: 2 hdry ¼ 1:35ðtm Þrm =rL þ K1 UV;valve rV =rL

(14.82a)

2 hdry ¼ 273:4K2 UV;valve ðrV = rL Þ

(14.82b)

(b) Units full open:

where tm is valve thickness in mm; rm is valve metal density, typically 8170 kg/m3 for ferrous valves, K1 , K2 are pressure drop coefficients; UV ,valve is velocity of gas through valve in m/s, as given by Eq. 14.84 and Avalve is total valve area per tray. Koch-Glitsch valve type V-1 is used for most applications. It has integral leg and is suitable for tray deck thickness up to 9.5 mm (3/800 ). The coefficients for this type of tray are K1 ¼ 54.686 (0.2 in fps units) and K2 as function of deck thickness is listed in Table 14.16.

Table 14.16 K2 as a function of deck thickness. Deck thickness mm (in.) K2

1.88 (0.074)

2.64 (0.104)

3.4 (0.134)

4.75 (0.187)

6.35 (0.25)

1.18

0.95

0.86

0.67

0.61

Alternatively, The expression of hdry in mm of liquid head is 2  hdry ¼ 273:4
K
UV;valve
ðrV = rL Þ

(14.83)

where UV;valve ¼

QV Avalve

(14.84)

14.3 Tray design

483

Avalve is the total valve area per tray in m2, QV is the vapour flow rate in m3/s, UV;valve is in m/s and r in kg/m3. The values of K for different tray deck thickness are given in Table 14.17.

Table 14.17 Dry tray pressure loss coefficient for different tray thickness. K

Tray deck thickness (mm) 2

1.05

3.25

0.82

5.8

0.58

Tray manufacturers also use equations of the form  3 hdry ¼ C1 þ C2 UV;valve =2gc where C1 and C2 are proprietary constants. Total tray pressure drop (htray , mm of liquid) is calculated from the following equation htray ¼ hdry þ 554
ðQL =lw Þ2=3 þ 0:4hw

(14.85)

where hdry calculated from Eq. 14.82 is in mm, hw is weir height in mm, lw is weir length in m, QL is in m3/s. Downcomer pressure drop hdc;prdrop in mm of liquid head loss for liquid outflowing from downcomer is given by 2  hdc;prdrop ¼ 177:7 UL;dc (14.86)   UL;dc ¼ QL A is the velocity of liquid outflow in m/s from under the downcomer onto dc;clarance the tray, Adc;clarance ¼ hdc;clarance
lw as discussed earlier and hdc;clearance is downcomer apron clearance, usually 40 mm for a 50 mm weir height. 16. Downcomer backup Downcomer backup (hL;dc mm) expressed as Eq. 14.87 shall not exceed 40% TS for high vapour density systems  48 kg/m3 (3 lb/ft3), 50% for medium vapour density systems, 16e48 kg/m3 (1e3 lb/ft3) and 60% for low vapour density systems 1.5 m Pitch: rotated square and triangular pattern, with pitch minimum twice the hole diameter; normal range is 2.5e4 times do Maximum number of valves: 130 to 150 per square metre of active area Mechanical aspects The following are the recommendations from the M/s Koch-Glitsch Manual. Distance of valve from weir and truss line A gap of 108 mm (41/400 ) is kept between (i) centreline of nearest valve from outlet weir, (ii) centreline of nearest valve from inlet edge of tray and (iii) between valve centreline and truss line for lap joints. The distance is 83 mm (31/400 ) for butt joints. All distances may be changed in case of special applications. Distance of valve centreline distance from tray ring should be minimum 32 mm (11/400 ). Tower manhole ID If the number of valves per panel is 5, 6 or 7, using 63.5 mm (21/200 ) as distance between row centres, the manhole approximate ID is 406 mm (1600 ), 470 mm (181/200 ) and 533.4 mm (2100 ), respectively. Large manhole diameters are important for large diameter towers as this can substantially reduce the number of panels.

14.4 Packed tower

485

Trusses and beam(s) Small towers have trusses parallel to liquid flow. Major support beam(s) may be required if tray diameter is more than 3.6 m (12 ft). The beams are placed parallel to flow and the trusses are perpendicular to it. Truss depth and construction is based on weight of tray plus uniform load of 97.7e122 kg/m2 (20e25 lb/ft2). Maximum allowable tray deflection is 3 mm (1/800 ) for towers up to 3.8 m (12 ft 6 in.) and 4.8 mm (3/600 ) for higher diameters. Trusses in addition are designed to bear a point load of 113.5 kg (250 lb) on any point without exceeding tangential stress limit at extreme fibres. Explosion proof trays are designed for 2930 kg/m2 (600 lb/ft2) loading. Circular downpipes and rectangular ducts These are used at transition trays, chimney trays, accumulator trays and from the bottom tray sump to tower bottom. The collection area or the recessed sump is sized with the same considerations as the downcomer. A sump should be minimum 380 mm (1500 ) deep. The duct velocity is kept 0.6e0.9 m/s (2e3 ft/s). Table 14.19 Standard dimensions on trays for different pressure services. Service

Vacuum

Atmospheric

High pressure

TS(m)

0.4e0.6

0.4e0.6

0.3e0.4

lw (m)

0:5  0:6D

0:6  0:76D

0:85D

hw (m)

0.02e0.03

0.03e0.08

0.04e0.1

Skirt clearance hdc;clearance (m)

0:7hw

0:8hw

0:9hw

Bubble cap diameter dcap (m)

0.08e0.15

0.08e0.15

0.08e0.15

Bubble cap pitch

1:25dcap

ð1:25 1:4Þdcap

1:5dcap 0.04e0.05

Valve diameter dvalve (m)

0.04e0.05

0.04e0.05

Valve spacing (m)

1:5dvalve

ð1:7 2:2Þdvalve

Hole diameter do (m)

0.004e0.013

0.004e0.013

0.004e0.013

Hole spacing

ð2:5 3Þdo

ð3 4Þdo

ð3:5 4:5Þdo

Relative free area (%)

15e10

10e6

7.5e4.5

14.4 Packed tower Packed towers are used as gaseliquid and liquideliquid contacting equipment. In most applications, the fluids flow in countercurrent fashion. Mass transfer applications using gaseliquid flow in packed bed are distillation, absorption and stripping and liquideliquid flow is involved in extraction. Adsorption and desorption in packed bed involves a single flowing phase (gas/liquid) through a packed bed where solute transfer takes place from (to) the fluid bulk to (from) the packed bed of adsorbent. Apart from the mass transfer applications, packed beds are also used in regenerative heat exchangers and packed bed reactors. These cases mostly involve a single fluid flowing through the bed. This section focuses on design of packed bed for vapoureliquid contacting e primarily for distillation, absorption and stripping. A packed bed contactor would comprise of the shell containing

486

Chapter 14 Column and column internals for gas

packing in one or more sections, packing support(s), liquid distributor, liquid collector and gas and liquid entry and exit nozzles. Intermediate supports and redistributors are used in case of tall columns. General arrangements in a typical packed tower are presented in Fig. 14.20.

FIGURE 14.20 Schematic diagram of a packed tower.

Countercurrent flow of gas and liquid throughout the packing makes it more effective for mass transfer as compared to tray towers that involve cross flow of gas and liquid on each tray. The contacting, however, is not perfect due to the nonhomogeneous structure of packing causing nonuniform flow over the cross section. Maldistribution of liquid is usually caused by channelling and wall flow. Appropriate packing shape and size is selected to minimise this maldistribution while ensuring high capacity, flexibility in gas and liquid throughput combined with high but nearly constant separation efficiency. Packing elements are classified as • •

Random e these can be ‘stacked’ or ‘dumped’ to form the bed. Typical shapes are standard and many are proprietary. Fig. 14.21 shows several such random packing elements. Structured e these are typically blocks of suitably spaced corrugated plates. These blocks are arranged side by side as well as in layers and are stitched with wire to form the bed. Near-vertical alignment of different layers of vapour passage reduces form friction. The skin friction pressure drop is contributed by the interaction between gas and liquid film on the packing surface. The

14.4 Packed tower

487

proprietary shapes of the corrugations, inter plate protrusions and surface holes are employed to increase surface area and mass transfer. Compared to random packing, structured packing allows a higher rate of mass transfer and lower pressure drop for the same bed height. A typical section of structured packing from M/s Sulzer Ltd., Switzerland, is shown in Fig. 14.22.

FIGURE 14.21 Commercial random packing elements: (A) Raschig ring, (B) Lessing ring, (C) Partition ring, (D) Berl saddle, (F) Tellerete, (G) Pall ring.

FIGURE 14.22 Structured packing from M/s Sulzer. © Sulzer Ltd. 2019.

488

Chapter 14 Column and column internals for gas

Packing manufacturers have been evolving competitive designs to prove the superiority of their packing over available design in terms of one or more desirable features listed below. (a) Low weight per unit volume. This affects not only the total weight to be carried by the tower but also the design of the tower shell itself. A packing that is dumped into the tower at random may exert a side thrust against the walls, and if the packing has a high unit weight, this may affect the cost of tower construction. (b) Large active surface per unit volume (c) Large free cross section. This is of importance because it affects the frictional pressure drop through the tower and the power required to circulate the gas. Also, a lower free cross section means higher gas velocity for a given throughput that may lead to an earlier onset of flooding. (d) Large void volume permits passage of high flow rate of fluid through small tower cross section without loading or flooding and achieve low pressure drop for gas. It is also desirable that the gas pressure drop is predominantly due to skin and not form friction. This is important in cases like absorption of oxides of nitrogen where sufficient time must be allowed for reactions in the gas phase. (e) Low liquid holdup e This is generally an advantage since it decreases the load on the tower and removes the liquid from the tower as rapidly as possible. In some cases, it may be a disadvantage, especially where the reaction between gas and liquid is slow or where the solubility of gas in liquid is not reasonably high. In towers handling hazardous liquids, it is particularly desirable to have lower amount of liquid retained as it lowers the potential hazard. Ideally the packing needs to be irrigated with a thin layer of liquid. (f) Inexpensive, reasonable mechanical strength and chemically inert towards the components involved. Design of packed tower apart from the normal features of any tower involves -

Selection of packing e type (random/structured), shape, size and material Packed bed details e depth, diameter and number of beds for a particular service Arrangements for liquid distribution/redistribution and draw off, if warranted

14.4.1 Choice of packing Packing types and size Choice of packing affects bed depth. Raschig rings and Pall rings are most common random packings. Saddle shapes are also popular but more expensive. As outlined earlier, structured packings are more expensive but require lower bed depth and incur lower pressure drop. In case of random packing, the size needs to be compatible with the tower diameter as use of large packing size relative to tower diameter leads to poor contacting. Typically, for uniform liquid and gas distribution, the ratio of tower diameter to packing size should be 10:1.Small size packings tend to flood more and the liquid load has an upper limit. Typical recommended values for minimum tower diameter and maximum liquid loading are presented in Table 14.20.

14.4 Packed tower

489

Table 14.20 Minimum tower diameter and maximum liquid loading for random packing. Random packing nominal size, mm (in.)

Minimum column inside diameter, mm

Maximum liquid loading, m3/hr per m2

19(¾)

250

0.53

25 (1)

300

0.84

38 (1½)

450

1.16

50 (2)

600

1.48

1100

2.64

89 (3½)

Retrofit designs for augmenting capacity and/or separation performance are carried out by replacing some/all existing trays with packed (structured packing) section(s). This works because the ‘height equivalent to a theoretical plate’ (HETP) and the pressure drop are lower for structured packing. Almost all crude distillation towers in Indian refineries using tray have today augmented capacity and separation efficiency by replacing some of the trays with sections of structured packing. It may also be noted that modern atmospheric cooling towers (see Chapter 7) involving counterflow of water and air are designed with structured packing. Table 14.21 lists the characteristics of random packings required by the designer.

Table 14.21 Packing factors for random and structured packing.

Type

Material

Nominal size, mm (in)

ε

 ap m2 =m3 ft2 =ft3

Fp mL1 (ftL1)

Relative mass transfer coefficient

Random packing Raschig rings

Berl saddles

Ceramic

Ceramic

13 (0.5)

0.64

364 (111)

1900 (580)

1.52

25 (1)

0.74

190 (58)

587 (179)

1.2

38 (1.5)

0.73

121 (37)

312 (95)

1.0

50 (2)

0.74

92 (28)

213 (65)

0.85

13 (0.5)

0.62

466 (142)

787 (240)

1.58

25 (1)

0.68

249 (76)

361 (110)

1.36

105 (32)

148 (45)

50 (2) Pall rings

Metal

25 (1)

0.94

207 (63)

184 (56)

1.61

38 (1.5)

0.95

128 (39)

131 (40)

1.34

50 (2)

0.96

102 (31)

89 (27)

1.14 Continued

490

Chapter 14 Column and column internals for gas

Table 14.21 Packing factors for random and structured packing.dcont’d Fp mL1 (ftL1)

Relative mass transfer coefficient

Material

Nominal size, mm (in)

Metal Intalox IMTP

Metal

25 (1)

0.97

230 (70)

134 (41)

1.78

50 (2)

0.98

98 (30)

59 (18)

1.27

Nor-Pac

Plastic

25 (1)

0.92

180 (55)

82 (25)

50 (2)

0.94

102 (31)

39 (12)

25 (1)

0.96

177 (54)

148 (45)

1.51

50 (2)

0.97

95 (29)

85 (26)

1.07

25 (1)

0.92

180 (55)

82 (25)

50 (2)

0.94

102 (31)

39 (12)

0.95

249 (76) 499 (152)

66 (20) 112 (34)

Flexipac 2 4

0.93 0.98

223 (68)

72 (22) 20 (6)

Gempak 2A 4A

0.93 0.91

220 (67) 452 (138)

52 (16) 105 (32)

Norton Intalox 2T 3T

0.97 0.97

213 (65) 177 (54)

56 (17) 43 (13)

299 (91)

108 (33)

700 (213) 492 (150)

230 (70) 69 (21)

Type

Hy-Pak

Metal

Plastic

ε

 ap m2 =m3 ft2 =ft3

Structured packing Mellapak 250Y 500Y

Metal

Montz B300 Sulzer CY BX

Wire mesh

0.85 0.90

1.98 1.94

Geankoplis, C. J., (2003). Transport processes and separation process principles (unit operations) (4th ed.). Reprinted by permission of Pearson Education, Inc., New York, NY.

14.4.2 Liquid distribution Ideal operation for vapoureliquid contacting is with a thin liquid layer flowing on the packing surface while in contact with the flowing gas. The entire packing surface is wet and the liquid layer continuously gets renewed. It is important to ensure uniform liquid distribution across the bed section by maintaining the liquid flow rate above a minimum limit. Tall columns are often fitted with liquid redistributors as maldistribution sets in after a certain depth from the top. It is also important to remain away from flooding, the other limit of operation.

14.4 Packed tower

491

Liquid distributor Four basic types of distributors are as follows: (a) Pan type (riser tube) e These have tall riser tubes (typically 100 mm) allowing the vapour to bypass the distributor, while the liquid pool overflows through downcomers with 25 mm weir height projecting from the pan. Schematic of pan distributor is shown in Fig. 14.23. (b) Pipe orifice headers (gravity or pressure type) e These distributors are fixed with flanges to the liquid inlet nozzle. The branches are fitted inside the tower. Even though holes are preferred, slots are used as they provide more opening area. The openings point downwards and their number per square metre is kept as uniform as possible. (c) Trough distributors are used with the liquid overflowing from the weir wall of the trough. These are common in large diameter towers and in cooling towers. (d) Spray nozzle headers are used when the liquid is available at adequate pressure, is clean and the gas velocity is not too high to cause entrainment. Nozzle pressure drop is usually kept below 1 kg/cm2 to avoid generation and entrainment of fine droplets. Sprays are located considering the cone angle of the spray, details of which is provided by the vendor.

Distributor feed line 6 “NB, 40 sch, pipe

Blocked end 50 250

Flange for fixing with inlet nozzle close to vessel wall

4 “NB, 40 sch pipe welded

10 300

Liquid downcomer(s) splash baffle

450

150 φ 600 φ

12 φ weep hole

Distributor deck

25 2400 φ

FIGURE 14.23 Schematic of a typical pan type distributor.

Vapour riser(s) 4 “NB, 40 sch pipe welded

492

Chapter 14 Column and column internals for gas

Design guidelines e gravity-type distributors • • • • •

Drip points to be located uniformly over tower section. For D  920 mm, 75e150 mm square pitch and for D > 920 mm, number of points z (D/150)2. Maximum number of points: 105 nos./m2; up to 95 in many services with random and structured packing. Minimum opening size: 10 mmf for carbon steel, 3 mmf for alloy steel. Total hole/slot area is calculated based on pressure drop of 170e350 mm WC and discharge coefficient of 0.6. This area is then distributed in different branches, etc. Liquid distributed within a distance of 5%e10% of D from tower wall should be kept below 10% to avoid liquid flow towards the wall. Structured packings are more prone to initial maldistribution of liquid.

Redistributor and collector Redistributors are usually not required in case of stacked packing as these have low channelling tendency. Redistribution for random dumped bed is recommended after every three tower diameters (or approximately 3e5 m) of bed depth for Raschig rings and 5 to 10 diameters (5e6 m) for saddle packings. The gravity redistributor is similar to a chimney tray with sieves or with drip tubes and risers for redistribution. Its inclusion typically adds about 1500 mm to column height. Beds of high efficiency packing may employ a collector above the redistributor to ensure (A) free passage of gas and (B) mixing of liquid across the column section. A collector with a distributor increases the column height significantly e typically by around 2500 mm. Typically, chimney tray is used for side stream draw below packed section in distillation columns. This also facilitates vapour redistribution. Chimney height decides the liquid residence time on a tray. It should be sufficient to allow level variation within 150 mm on the tray and the minimum chimney height is 225 mm from the tray floor. Minimum 300 mm clearance is kept above the chimney tray hat up to the upper tray deck. Manholes are provided above the packed bed with preferred clearance of 1200 mm from the upper deck for a 600 mm manhole.

14.4.3 Bed support Packing supports must offer low pressure drop and the free area of support should be higher than the packing voidage and rarely below 65%. Typically, bed height per support plate is 3.7 m for Raschig ring and 4.6e6.2 m for other packing shapes. The supports could be simple grids or perforated plates with risers or other designs. These can also be fabricated from expanded metal sheets of adequate strength. Simple grids are usually used for smaller columns. Bigger diameter columns use profiled grids where the support is not in one plane. This allows greater free area per tower cross section and also allows liquid and gas to pass in a segregated fashion. A low design pressure drop (w8 mm WC) across the support plate is considered for most applications to ensure that no liquid accumulates on the support plate. Mechanical design of the support is based on the weight it has to carry. This includes (A) weight of packing, (B) flooded liquid volume in the voids, (C) force due to pressure surges, if considered, (D) weight of any redistributor, if the same is also to be supported.

14.4 Packed tower

493

14.4.4 Flooding and pressure drop in randomly packed bed Operating region of a gaseliquid contactor is the flow rate range within which the contactor retains its desirable mass transfer rate and operability. In trays, this region is fixed by the type of tray and its dimensional details. In case of packed bed, the performance also depends on liquid distributor and the effective operating region of a packed bed is very different from that of a contacting tray. At extremely low liquid rates, the packing does not remain well irrigated, resulting in a drastic fall in mass transfer rate. This minimum mass flux of liquid is termed as the minimum wetting rate (MWR). A similar sharp fall is also observed at high gas and liquid rates as the operation approaches ‘flooding’. Thus, the limits of operation of the packed bed are the approach to bed flooding at high gas and liquid flow rates and the MWR at low liquid flows. The liquid distributor is functionally efficient only over a limited range of liquid mass flux and its efficacy is practically independent of the gas rate. Thus the effective operating region for a typical bed is the common area for distributor and bed as marked in Fig. 14.24. It is clearly seen that the operating region is actually limited by the liquid distributor and there can be a wide range of variation of gas flow rate, whereas the range of liquid flow rate is rather narrow. This apparent limitation may be utilised by the designer as an advantage by employing packed bed contactors for applications where the range of gas flow is wide or is rather uncertain. 101

Superficial gas velocity, m/s - ->

Liquid distributor operating limits

Wetting limit 100 Flooding limit

10–1 10–4

10–3

10–2

10–1

Superficial liquid velocity, m/s - ->

FIGURE 14.24 Operating range of a packed column.

Bed diameter estimation based on flooding and pressure drop Packed towers are typically used for D  1 m (3.3 ft). Pressure drop in packed bed depends on the packing characteristics, bed diameter, gas and liquid flow rates and their properties. The generalised pressure drop curves for randomly packed tower are

494

Chapter 14 Column and column internals for gas

shown in Fig. 14.25A and that for structured packings in Fig. 14.25B. The packing factors (Fp ) are listed in Table 14.21. For using Fig. 14.25A and B, Fp in ft/s is noted from Table 14.21. y is liquidphase kinematic viscosity in centistokes, vG is superficial gas velocity in ft/s (1 m/s ¼ 3.281 ft/s), rG ; rL are gas and liquid densities in lb/ft3 (1 kg/m3 ¼ 0.06243 lb/ft3). Note that though the abscissa is dimensionless, the ordinate is not. Validity of the generalised correlations in Fig. 14.25 requires the maximum limit of irrigation rates to remain below the values mentioned in Table 14.22.

(A)

2.4

ΔP = (in. H2O/ft) 2.0 1.5 1.0

2.0 1.6

0.50

υG[ ρG /(ρL – ρ G)]0.5 FP 0.5ν 0.05, 1.2 capacity parameter

0.25

0.8

0.10 0.05

0.4 0.0

0.006 0.02 0.04 0.06 0.20 0.40 0.60 0.005 0.01 0.30 0.50 1.0 0.03 0.05 0.10

2.0

4.0 3.0 5.0

(GL/GG)(ρG /ρL)0.5, flow parameter

(B)

3.0 2.8 2.6 2.4 2.2 2.0 1.8 υG[ ρG /(ρL – ρG)]0.5 FP 0.5ν 0.05, 1.6 1.4 capacity parameter 1.2 1.0 0.8 0.6 0.4 0.2 0.0

ΔP = 2.0 (in. H2O/ft) 1.5

1.0 0.50 0.25 0.10 0.05

0.01

0.05

0.10

0.50

1.00

2.00

(GL/GG)(ρG /ρL)0.5, flow parameter

FIGURE 14.25 Generalised pressure drop curves for (A) random packing, (B) structured packing. From (A) Strigle, R.F., Jr. (1987). Random packings and packed towers: design and applications. Gulf Publishing Company; (B) Kister, H.Z. (1992). Distillation design. New York: McGraw-Hill Book Company.

14.4 Packed tower

495

Table 14.22 Upper limit of validity of irrigation rate (liquid viscosity below 3 cP) for random packing. Upper limit of irrigation rate, m3/hr per m2 tower cross section

Nominal packing size, cm (in.) 1.6 (5/8)

41.5

2.5 (1)

95.5

1

3.8 (1 /2)

134.5

5 (2)

166

Each curve in Fig. 14.25 represents a value of constant pressure drop per unit bed depth as marked thereon. The empirical correlation to estimate pressure gradient at flooding is for Fp  197 m1 (60 ft1) DPflood ðin: = ftÞ ¼ 0:115
ðFp Þ0:7 and when Fp > 197 m

1

(14.89a)

(60 ft1) DPflood ¼ 2 in:=ft

(14.89b)

This allows the flooding line to be located as shown in Fig. 14.25A or B for specific packing type and size and the design vapour velocity is calculated corresponding to j % approach to flooding. The tower area and diameter are calculated from the gas flow rate and the design vapour velocity. It is preferred to operate the packed bed close to loading conditions. The onset of loading is usually around 65%e70% of flooding velocity. For absorption, the tower is designed for 50%e70% of gas flooding velocity with the higher value adopted for higher flow parameters. For atmospheric pressure distillation, the value is 70%e80%. For distillation and structured packing, 80% of flooding velocity is normally used in design. For extraction process, where both the phases are liquid, the flooding velocity can be predicted from Crawford and Wilke correlation. This along with the approach to flooding and packed column characteristics for extraction process is discussed in Chapter 13.

Pressure gradient Pressure drop per unit bed length is noted from the constant pressure drop curve passing through the intersection of the design X and Y value in Fig. 14.25. Typical recommended values for design pressure drop for randomly packed beds in several services are shown in Table 14.23.

496

Chapter 14 Column and column internals for gas

Table 14.23 Typical design pressure drop for random packing. Pressure drop (mm WC/m bed depth)

Application Absorber/regenerator e nonfoaming service

20 to 35

Absorber/regenerator e moderately foaming service

12 to 20

Fume scrubbers e water absorbent

35 to 50

Fume scrubbers e chemical absorbent

20 to 35

Fractionating towers (close to atmospheric or higher pressure)

35 to 100

Vacuum towers

12 to 35

Minimum wetting rate The bed design must ensure that the liquid flux is above the MWR. As can be seen from Table 14.24, MWR largely depends on the packing material. Table 14.24 Minimum wetting rate for different packing material. Packing material

Minimum wetting rate (MWR) m3/hr per m2 packing surface

Unglazed ceramic

0.5

Oxidised metal (carbon steel, copper)

0.7

Surface treated metal (etched stainless steel)

1.0

Glazed ceramic

2.0

Glass

2.5

Bright metal

3.0

PVC

3.5

Polypropylene

4.0

PTFE

5.0

14.5 Packed tower design Process design of a packed tower involves choosing the packing type (shape), material and size; estimation of packed section diameter and height; location and space required for liquid distributor/ redistributor and draw-off as applicable. A layer (100e120 mm) of larger packing and/or alumina balls is sometimes placed over the active bed. This helps uniform liquid irrigation, acts partly as bed restrainer to prevent packing carryover with vapour and clogging of nozzle.

14.5 Packed tower design

497

This is followed by design of the mechanical details of packed section, packing support, bed restrainer and other fittings. Inputs for the design are • • • •

Liquid and gas inflow rates, mV and mL in kg/s Inlet compositions of both streams and desired change in concentration of transferred component(s) Maximum allowable pressure drop for each phase Properties of phases, rV ; rL , densities of vapour and liquid in (kg/m3), mL , liquid viscosity in cP (and possibly their estimation method based on composition, temperature and pressure) Outputs from the process design are

• •

Active bed height: Governed by the targeted change in concentration of component(s). Bed diameter: Decided from hydraulic considerations e flooding/loading and pressure drop in the tower.

In packed bed design, tower diameter and active bed height are closely coupled. In any specific design for stated inlet flow rates, choosing a lower bed diameter results in higher superficial velocity and operation closer to flooding/loading. A higher superficial velocity results in higher volumetric mass transfer coefficient ðkL aÞ and lower HETP. Thus, choosing a lower diameter subject to meeting the limit of flooding and gas stream pressure drop also results in lower bed depth and a shorter tower. The design methodology is as follows: (i) Selection of packing type (shape), material and nominal size e based on the considerations already outlined in Section 14.4.1. (ii) Rough estimate of bed diameter based on flow rates using information available in Table 14.20. (iii) Active bed height hbed ¼ N
HETP; where N is the number of theoretical plates and HETP is the height equivalent to a theoretical plate, calculated for the chosen packing characteristics based on mass transfer correlations. The procedure is already outlined in Chapter 10. One may also refer to design illustration P10.1 that illustrates the HTU-NTU method as an alternative approach. (iv) Tower height Once the bed height is estimated, the height required for distributors, bed supports, etc., are estimated conservatively. The summation of all heights gives the total height of the column. Information in Section 14.4.2 for distributors, redistributors and draw off is to be used for arriving at the total tower height. (v) Column diameter and pressure drop. Column diameter is estimated by using the generalised pressure drop correlation plot in Fig. 14.25. The procedure is outlined in Section 14.4.4. The design conditions plotted in Fig. 14.25 gives the pressure gradient in the active bed. Total pressure drop in the tower is the summation of active bed pressure drop and the additional drop across the following components

498

Chapter 14 Column and column internals for gas

-

inlet and exit nozzles bed support (plates, etc.) bed top restrainers or bigger size packing above it, if provided redistributors/draw off trays, if provided.

Increase in the ratio of packing size to tower diameter increases the tendency of liquid to channel towards the tower wall. Based on this, the minimum tower diameter suggested for different nominal size of packing is listed in Table 14.20. Packing installation During installation of random packings, the packings are poured into the tower randomly and in order to prevent breakage of ceramic or carbon packings, the tower is usually filled with water to reduce velocity of fall. A chute may be lowered to direct the packings at different locations in a cross section in case of large tower diameter if the filling is done in dry condition. The dry packed tower is denser and the pressure drop is higher (by about 50%e60%) as compared to wet packing. This is not the preferred method due to significant settling.

14.6 Chimney tray, reflux entry, feed tray and tower bottom 14.6.1 Chimney tray A good chimney tray design is essential in packed towers, especially in vacuum towers where pressure drop is critical, in order to provide uniform vapour distribution while minimising pressure drop. These have rectangular or circular chimneys with hat, located as uniformly as possible over the tray deck. Rectangular chimneys are more economic for large tower diameter. Providing a single central chimney is avoided as it leads to vapour funnelling. Rectangular chimneys are oriented with their longer side parallel to the liquid flow direction. This is done to minimise the resistance to liquid flow. Level on the tray and chimney height: Minimum chimney height is 200 mm. Highest liquid level on the chimney tray should be at least 100e150 mm below the chimney top edge. In order to get the level, liquid gradient on the tray is calculated from (a) liquid crest over weir (Francis’ weir equation e Eq. 14.14), (b) head loss during flow in between the chimney or the chimney and an obstacle (considering open channel flow) and (c) gradient for open channel flow (Manning equation). The value taken is the higher of (a) þ(c) or (b) þ (c). Typical chimney width can be 300e400 mm. With circular or square chimneys, the diameter or side is usually 450 mm. Chimney area: Chimneys on trays are provided with a minimum area which is larger of (a) 15% of tower cross section and (b) 1.5 times open area of the tray above. For chimneys below packed bed or grid, chimney area is based on maximum 6 mm WC pressure drop that usually leads to chimney area of 15%e30% of tower area. In case of vacuum services, the area provided is higher (30%e35%) in order to minimise pressure drop. Vapour flow area below hat: The area provided is equal to the chimney cross-sectional area. Spacing below the chimney tray: When the chimney tray is located above a tray, the spacing is higher of (a) 600 mm and (b) 150 mm þ tray spacing. For chimney trays above a grid or a packed bed, sufficient space to access the liquid distributor needs to be kept. Spacing above the chimney tray: Higher of (a) 300 mm, (b) width of chimney in case of rectangular chimneys, or 1.5 times hat diameter.

14.6 Chimney tray, reflux entry, feed tray and tower bottom

499

Chimney hat: Hat dimension should be 15% more than the chimney opening. Rectangular chimneys are provided with V-hat whose three sides should be upturned by w40 mm to collect the showering liquid and drain from one side. This reduces entrainment. Fig. 14.26 shows the details of a rectangular chimney with V-hat. In case of circular chimneys, maximum of three numbers of 75 mm wide notches are provided. Draw off box: Downward velocity in the draw off box should be limited to 0.3 m/s. Depth of the box, usually square, should be higher of (a) 150 mm and (b) 1.5 times the draw off line diameter. In larger towers, the box should be at least 450 mm
450 mm square to permit easy access. Nozzles flush to the bottom of the box (draw off nozzles) permit drainage during shutdown. In draw offs below packed beds, the liquid has to cross the tray to enter the draw off box. For large liquid rates, the chimney riser and weir height needs to be adjusted considering the liquid height. Hat width Chimney width

Upturned lip (3 sides) Vapour exit Hat

Chimney length

Chimney

Pictorial view

Top view

Hat lip Area for vapour exit from side

min. 140°

Area for vapour exit from side

Chimney height

Chimney width Narrow end view

Chimney length Long end view

FIGURE 14.26 Rectangular chimney schematic with V-hat.

500

Chapter 14 Column and column internals for gas

14.6.2 Reflux entry arrangement on top tray Reflux entering the top tray may be saturated or sub-cooled. Entrainment of the entering liquid is undesirable. In addition, any hydraulic hammer and vibration due to sudden contraction of the hotter rising vapour on contacting the subcooled reflux needs to be avoided. The liquid entering arrangement therefore contains a false downcomer and a horizontal baffle plate above the inlet, as shown in Fig. 14.27 for a single-pass top tray. The recommended minimum dimensions/clearances are also shown. For a two-pass tray, a perforated pipe distributor or a T-shaped inlet is used. Limiting dimensions of features a, b and c in Fig. 14.27 are as follows: (a) Minimum of 25 mm should be the downcomer bottom clearance (b) Minimum 100 mm (c) Normal tray spacing or 100 mm plus 1.5
manhole diameter, if a manhole is provided above the top tray. In case a demister is provided, the dimensions are w.r.t. to the elevation of the same.

Open top A Baffle

Baffle

A' (b) min.

(c) False downcomer (a)

Reflux entry nozzle Top tray

3 dN False downcomer

dN Open top

Section A-A'

FIGURE 14.27 Reflux entry arrangement on single-pass top tray with vapour nozzle on tower dome.

14.6.3 Feed tray Column feed may be vapour or liquid or a mixture of both. Feed location can be below the bottom most tray or in between trays. Liquid feed stream is normally introduced into downcomer leading to feed plate. This is not adopted for feed at bubble point or a liquidevapour feed as feed flashing on entry to tower may lead to downcomer flooding. There are some variations for the (feed) inlet arrangement in different cases. Perforated pipe distributors are more common for liquid and liquidevapour mixtures. Flushed nozzles may be used for vapour inlet. Tangential entry and chimney trays are used when the system is highly foaming. The most common type of distributor is the perforated pipe distributor, particularly when the pressure drop is affordable and there is limited vapour fraction in the inlet stream. Typical feed inlet arrangement on a single-pass tray with a straight perforated pipe distributor is shown in Fig. 14.28. The distributor arrangement should by symmetrical in case of multipass trays.

14.6 Chimney tray, reflux entry, feed tray and tower bottom

501

Discharge from the perforations is directed towards an insulating baffle (not shown in the figure)/ downcomer apron at an angle (45 for vertical downcomers and at a lesser angle for sloped ones). Round perforations usually provide more uniform distribution but lesser area compared to slots. Single straight distributors are used for tray tower diameter up to 3000 mm. Branched distributors (T, U, H or multiple leg) are used in large towers where distribution uniformity is critical.

A'

A (b)

(c)

(a) 45° (d)

A - A' section

FIGURE 14.28 Feed inlet arrangement on a single-pass tray with a perforated pipe distributor.

Limits on the dimensions/features marked on the figure are as follows: (a) (b) (c) (d)

At least equal to tray spacing (higher of that above and below feed tray). Minimum 75 mm. Minimum 200 mm. 45 for vertical downcomers and a lesser angle for sloped ones.

Insulating baffle may be installed at a horizontal distance of 12e25 mm from downcomer if feed temperature is above temperature of downcomer liquid. Insulation baffle prevents the hot feed impinging on the downcomer apron that may lead to vaporisation in the downcomer. Perforated distributor: The distributor having the same diameter as the inlet nozzle is connected to it with a flange inside the tower. Due to this, there are no perforations on the first 300 mm of the distributor or 10% of distributor length, whichever is higher. Minimum gap between perforations is the thickness of the pipe. Minimum perforation size is 13 mm. Furthest end of the distributor is blanked off and provided with a 6 mm diameter drain hole to empty the distributor when not in use. The perforations are sized based on pressure drop of 13e25 mm Hg at design flow rate. Though higher pressure drop improves distribution, exceeding the maximum limit may lead to premature flooding of the feed tray. Equation used for estimating the pressure drop is based on mixture velocity, mixture density and orifice coefficient of 0.6 e all of these can be found from weight fraction of vapour at inlet, feed mass flow rate and density of the liquid and vapour. A 13 mm drain hole is provided at the downstream end of each distributor if the process fluid does not discharge vertically downwards.

502

Chapter 14 Column and column internals for gas

Flashbox design: In case of a large proportion of vapour in inlet, the perforated pipe distributor has perforations on both sides and is enclosed between two baffles as shown in Fig. 14.29. The minimum flashbox width is higher of (a) 200 mm plus pipe diameter and (b) twice pipe diameter. Inlet velocity is limited to 30 m/s and the slot velocity limit is 2 m/s. Flashbox top to upper tray distance is the higher of (a) 450 mm and (b) 70% of tray spacing. Its depth should be sufficient to ensure that the exiting fluid falls within the box itself.

(A)

Splash baffles

Feed inlet Distributor (perforated)

Closed end

(B)

Slots

No slot zone

Feed inlet

Slots

Closed end

FIGURE 14.29 Typical flashbox arrangement for feed inlet: (A) side view, (B) top view.

14.6 Chimney tray, reflux entry, feed tray and tower bottom

503

14.6.4 Tower bottom arrangement Tower bottom will have arrangement of inlet for vapour in case of strippers, absorbers, adsorbers (gasesolid), humidifiers and kettle-type reboiler. A flush nozzle can be used below the bottom tray or between trays when vapour added is 6% of tray vapour load and distribution is not critical. For all other cases, a distributor should be used. The inlet stream is a mixture of vapour and liquid in case of thermosyphon reboilers and forced circulation reboilers. Since the column internals are affected by reboiler type, a brief discussion is presented below to decide on the reboiler type to be employed, if not already decided. e e e

Thermosyphon reboiler: This is the most common type with once through passage, returning a vapoureliquid mixture with maximum 80% vapour. Kettle-type reboiler: This is used when the boiling range is narrow and a vapour stream is returned to the column bottom. Recirculation-type reboiler: This is fed with liquid from the column bottom and the reboiler exit is returned to the tower. High recirculation rate achieves low vapour fraction at reboiler exit. Usually a single nozzle is used for withdrawing the bottom product and the reboiler feed.

A simple arrangement of inlet, draw off and a bottom tray downcomer with seal pan is shown in Fig. 14.30.

Seal pan

Reboiler return

•

Bottom tray downcomer

Bottom tray Reboiler return min 150

TS + 150 •

min 150 Liquid residence

Vortex breaker

min 150 BTL

FIGURE 14.30 Typical arrangement at column bottom.

504

Chapter 14 Column and column internals for gas

14.7 Design illustration Process design of a fractionator for recovering benzene from a mixture of benzene, toluene and orthoxylene is presented as design illustration in Section 11.5. This section considers two identical towers with 50% capacity each and details the tower internals. The flow conditions as specified in the problem are as follows: QF ¼ 20 m3/hr; rF ¼ 8755 kg/m3; mF ¼ 20
8755 ¼ 175100 kg=hr ¼ 48:64 kg=s; mD ¼ 495:8
78 þ 26:10
92 ¼ 41073:6 kg=hr ¼ 11.41 kg=s; mB ¼ 165:3
78 þ 446:1
92 þ 755:6
106 ¼ 134028:2 kg=hr ¼ 37.23 kg=s. Tower height The number of trays as estimated in the problem is N ¼ 5ðrectification sectionÞ þ 2ðstripping sectionÞ þ 1ðfeed trayÞ ¼ 8 Assuming TS ¼ 600 mm, from discussions on tray spacing in Section 14.3 the column height is obtained as 7(600) þ900 (1.5 TS over feed tray) þ 750 (150 mm þ TS for bottom downcomer) þ 1200 ¼ 7050 mmz 7600 mm. The scale of operation suggests that either bubble cap or valve tray tower can be suitable for this ternary fractionation problem. The tray design is illustrated for both tray types. Illustration Problem 14.1 considers bubble cap tray. Ideally for an accurate design, the top tray, bottom tray and feed tray should be designed separately and the largest diameter obtained should be adopted as the tower diameter. However, in this case only the complete design of the top tray is worked out for a distillate flow rate of 20532.2 kg/hr (50% capacity) for average molecular weight of mwD ¼ 78.7 and reflux ratio of 2.336. The reader is advised to repeat the same for the other two cases to arrive at an optimised design. Problem 14.1.

Design of a bubble cap tray for the top tray condition specified as mV ¼ DðR þ 1Þ ¼ ð2:336 þ 1Þ20532:2 ¼ 68495:42 kg=hr ¼ 19.026 kg=s mL ¼ RD ¼ 2:336
20532:2 ¼ 47963:22 kg=hr ¼ 13.323 kg=s rV ¼

78:2
273
164:2 ¼ 4:18283 kg=m3 22:4
ð273 þ 98:6Þ
101:325

rL ¼

4:18283 ¼ 878:2 kg=m3 19336:2=878:7 þ 1196=863:6

QL ¼

47963:22 ¼ 54:6 m3 =hr z 0.015171 m3 =s 878:2

QV ¼ 16375:3774 m3 =hr ¼ 4:55 m3 =s

14.7 Design illustration

505

Tray diameter Estimation by Fair and Mathews qffiffiffiffi Correlation rV mL From Eq. 14.6, FLV ¼ m rL ¼ 0.048; lnðFLV Þ ¼ ð 3:03655Þ V Using Eq. 14.8, lnðCsb Þ ¼ a3flnðFLV Þg3 þ a2flnðFLV Þg2 þ a1flnðFLV Þg þ a0  1.188 ¼ ð2.331152648Þ i.e., Csb ¼ 0.0972 qffiffiffiffiffiffiffiffiffiffiffiffiffi  V ; UfV;n ¼ 1.4 m s and from Eq. 14.9, for j ¼ 70% approach to From Eq. 14.7, Csb ¼ UfV;n ðr rr Þ L V flooding sffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi !   0:2  j ðrL  rV Þ 23 0:2 UV;n ¼ ¼ 1:011 m=s ¼ 1:4
0:7
ðs=20Þ Csb 100 rV 20 This gives tower diameter from Eq. 14.10 as follows for downcomer area fraction, k ¼ 12%


1=2 qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi 4ðmV =rV Þ 19:026 D ¼ pð1k=100Þ U1V;n ¼ p4
1:011
4:18283
0:88 ¼ 2:55 mz2:6 m on rounding off based on Table 17.8 and A ¼ 5:31 m2 Estimation by Souders and Brown Correlation For a nonfoaming system, Kcorr ¼ 1 from Table 14.3b and a ¼ 119:59 and b ¼ 285:6 from Table 14.3a. for TS ¼ 600 mm. From Eq. 14.5, C ¼ 0:3048
Kcorr fa lnðsÞ þbg ¼ 0:3048
119:59 lnð23Þ þ 285:6 ¼ 201:343 This gives mV;max from Eq. 14.4 as 0:5 mV;max ¼ CfrV ðrL  rV Þg1=2  2  4:18283Þ  2¼ 201:343½4:18283ð878:2 ¼ 12173:95311 kg= m hr ¼ 3.38165 kg= m s

Tray diameter D is obtained from Eq. 14.3 as


  ffi 1=2 sffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi 4 mV 4 19:016 D¼ ¼ ¼ 2:67 m p mV;max p 3:38165

For the rest of the calculations we take D ¼ 2:6 m. Check for entrainment From Fig. 14.11, for FLV ¼ 0:048, j ¼ 70%, j ¼ 0:045 < 0:1 (Ok) This gives total amount entrained e (in kg mol/hr) for liquid flow rate NL ¼ (47,963.22/78.1) ¼  jNL ¼ 28.6 kg mol hr. 609.4437 kg mol/hr from Eq. 14.11 as e ¼ 1j The decreased tray efficiency Eactual due to entrainment for Emv ¼ 0:7 from Eq. 14.12 is Eactual ¼ Emv

1 ¼ 1 þ eEmv =NL

0:7 ¼ 0:68 28:6
0:7 1þ 609:4437

506

Chapter 14 Column and column internals for gas

Number of passes  From Table 14.5, for QL ¼ 54:6 m3 hr and D ¼ 2:6 m, we select single-pass cross-flow tray and consider lw ¼ 0:76
D ¼ 1976 mm; Liquid loading QL =lw ¼ 54:6=1:976 ¼ 27:63 m3/hr (m lw ) 1:2 m, cap nominal size, dcap ¼ 100 mm (400 ). From Table 14.6, we select carbon steel caps of 100 mm nominal size for which Cap OD ¼ 104 mm, ID ¼ 98:4 mm; height ¼ 76 mm, p Cap area ¼ Acap ¼ ð0:104Þ2 ¼ 8:495
103 m2 4 Height of shroud ring ¼ 6 mm; skirt height ¼ 25 mm. Trapezoidal slot with number of slots ¼ 26. Rs ¼ 0:5, slot dimensions 8:5 mm
4:2 mm and hs ¼ 32 mm. Area ratios: reversal/riser ¼ 1.52; slot/riser ¼ 1.69; annular/riser ¼ 1.25; slot/cap ¼ 0.62; pitch P ¼ 1:25do ¼ 1:25
104 ¼ 130 mm; 6 pffiffi Considering triangular pitch, number of caps per m2 of net perforated area of tray Ao;tray ¼ 2
10 ¼ 3 P2 6 pffiffi2
10 2 ¼ 68:3z69 and 3
ð130Þ

Total number of caps ¼ 69
Ao;tray ¼ 69
3:2544 ¼ 224:55 z 224 Weir height From Eq. 14.13 and Table 14.6, for dcap ¼ 100 mm, hw ¼ cap skirt clearance (25 mm) þ shroud ring height (6 mm)þ slot height (32 mm) þ static seal (10 mm) ¼ 73 mm hw z 12%ðOKÞ TS Height over weir From Eq. 14.14, how (in mm liquid) for a segmental downcomer is   0:6667  mL 2=3 13:323 how ¼ 750 ¼ 750
878:2
1:976 rL lw ¼ 29:2 mmð > 10 mm specified as minimum in textÞ

14.7 Design illustration

507

Tray pressure drop From Table 14.6, Aslot =Acap Ariser 0:62 ¼ 0:367 ¼ ¼ 1:69 Aslot =Ariser Acap Ariser = cap ¼ 8:495
103
0:367 ¼ 3:11646
103 m2 Total riser area per tray in m2, Ar ¼ 3:11646
103
224 ¼ 0:6981 m2 . For r ¼ 1:25;Kc (from Eq. 14.25 for round bubble caps) ¼ 0.6373(1.25)22.0386(1.25)þ2.0554 ¼ 0.503. rV 4:18283 ¼ 0:004763 ¼ 878:2 rL From Eq. 14.24, hdry ¼ 273:4Kc

rV rL

   2 QV 4:55 2 ¼ 273:4
0:503
0:004763
¼ 27:825 mm 0:6981 Ar

Head loss through wet slots From Table 14.6, As =Acap ¼ 0:62 Slot area per tray, As ¼ 8:495
103
0:62
224 ¼ 1:18 m2 , hs ¼ 32 mm and Cs ¼ 0:74 for a trapezoidal slot with bottom width nearly twice the top width. qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi Substituting the values in Eq. 14.26. 1 QV;max ¼ 0:060478Cs As fhs ðrL  rV Þ=rL g2 ¼ 0:060478
0:74
1:18
32
ð878:24:18283Þ 4:18283 ¼ 4:3183 . 4:55 ¼ 1:05z100% and h ¼ h ¼ 32 þ 6 ¼ 38 mm (for 6 mm height of rQV QV;max ¼ 4:3183 so s shroud ring) Mean tray diameter ¼ ðD þ lw Þ=2 ¼ ð2600 þ 1976Þ=2 ¼ 2288 mm QL = ðMean tray diaÞ ¼ 0:015667=2:28 ¼ 0:006629

m3 =s ðOkÞ m

Tray Layout Based on tray diameter of 2600 mm, weir width of 455 and 100 mm calming zone. qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi Distance of weir from tray centre ¼ ð2600=2Þ2  ð1976=2Þ2 ¼ 845 mm qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi Width of calming zone ¼ 2
ð2600=2Þ2  ð1300  455  100Þ2 ¼ 2130 mm Average tray width where caps are placed ¼ ð2600 þ 2130Þ=2 ¼ 2365 mm Hence, average no of caps per row ¼ 2365/130 ¼ 18.19. No. of rows to accommodate 224 caps ¼ 224/18.19 ¼ 12.3, say 13  Gap between rows ¼ 130
sin 60 ¼ 112:6 mm

508

Chapter 14 Column and column internals for gas

Maximum number of rows that can be accommodated ¼ 2
ð845 100Þ=112:6 ¼ 13:23 > 12.3; hence, ok. Actual layout needs to be done to finalise number of rows but 12 rows can be accommodated and further calculations consider 12 rows. For the chosen cap (OD 104 mm, 400 ) and triangular pitch (130 mm). g ¼ (gap between caps/cap OD) ¼ (130  104)/104 ¼ 0.25; s ¼ Cap skirt clearance ¼ 25.4 mm. hlo ¼ hw þ how ¼ 73 þ 29:2 ¼ 102:2 mm In case of square pitch, number of caps per m2 of net perforated area of tray Ao;tray ¼

106 106 ¼ ¼ 59:1716 z 60 P2 ð130Þ2

and total number of caps ¼ 60
Ao;tray ¼ 60
3:2544 ¼ 195:264 z 196 Total riser area per tray in m2, Ar ¼ 3:11646
103
196 ¼ 0:611 m2     r QV 2 4:55 2 hdry ¼ 273:4Kc V ¼ 273:4
0:503
0:004763
¼ 36:32 mm 0:611 r L Ar Slot area per tray, As ¼ 8:495
103
0:62
196 ¼ 1:0323 m2 QV;max ¼ 0:060478
0:74
qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi . 32
ð878:24:18283Þ ¼ 3:778 m3 s which suggests overloaded slots. 1:0323
4:18283 Liquid gradient Factor f,  f ¼ 4831.18
m3 =s liquid flow per m of mean tray width ¼ 4831:18
0:006629 ¼ 32:025 From Eq. 14.21


ðQ=Lw Þ ¼ exp 0:0899
flnðf Þg2  0:0238
lnðf Þ þ 2:4146 ¼ 30.31 Cd And from Eq. 14.20

    ðQ=Lw Þ g 0:3
s 0:5
d
1:6
d þ 3
hw þ how þ D = 2 þ ¼ 0:2015
Cd 1þg g Combining the above two equation where D and d are unknowns and from Eq. 14.22, D ¼ Cv
No: of rows
d ¼ Cv
12
d Cv ¼ 1:16 is found from Fig. 14.13 against liquid load (gpm per ft of average tray width) with pffiffiffiffi ffi Vo rv as a parameter (Vo in ft/s and rv in lb/ft3). Relevant conversion factors: 1m3/ s ¼ 15850.32314 US gpm; 1m ¼ 3.281 ft; 1m/s ¼ 3.281 ft; 1 kg/m3 ¼ 0. 06423 lb/ft3.

14.7 Design illustration

509

This results in an implicit equation in d and its iterative solution yields d ¼ 2.9 mm/row and D ¼ Cv
no: of rows
d ¼ 1:16
12
2:9 ¼ 40:4, say 40 mm. Drop through aerated liquid From Table 14.6, riser slot seal hss ¼ 13 mm, how ¼ 29:2 mm; D ¼ 40 mm From Eq. 14.29, for how ¼ 29:2 mm, hds ¼ hss þ how þ D=2 ¼ 13 þ 29:2 þ 20 ¼ 62:2 mm 1=2

FVa ¼ UVa rV ¼

4:55 pffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi 4:18283 ¼ 2:306 5:31ð0:76Þ

From Eq. 14.28, b ¼  0:0255
ð2:306Þ3 þ 0:1744
ð2:306Þ2  0:4282
2:306 þ 0:9979 ¼ 0.6252 From Eq. 14.27, hal ¼ bhds ¼ 0:6252
62 ¼ 38:76 mm From Eq. 14.23, rhtray ¼ hdry þ hso þ hal ¼ 27:8 þ 38 þ 38:76 ¼ 104:56 mm Vapour distribution h ¼ hdry þ hso ¼ 27:8 þ 38 z 66 mm . cap From Eq. 14.30, rvd ¼ D hcap ¼ 40 66 ¼ 0:6 Corrected ‘approach to flooding’ From Table 14.8, An ¼ 0:88
5:31 ¼ 4:673 m2  4:55 ¼ 0:9737 m s From Eq. 14.31, UV;n ¼ QV An ¼ 4:673 ffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi q qffiffiffiffiffiffiffiffiffiffi  From Eq. 14.7, Uf V;n ¼ Csb rLrrV ¼ 0:0972 878:24:18283 ¼ 1:405 m s 4:18283 V

V ¼ j ¼ UUf V;n

0:9737 1:405

¼ 0:693z69:3% approach to flooding (Ok)

Downcomer dynamics From Eq. 14.16, wdc ¼ 0:455 m; From Table 14.8, Adc ¼ 0:12
5:31 ¼ 0:6372 m2 hdc;clearance ¼ 25 mm 2 From Eq. 14.19, Adc;clearance ¼ hdc;clearance
lw ¼ 0:025
1:976 ¼   20:0494 m 2 QL 54:6 From Eq. 14.32, hdc;prdrop ¼ 0:1275 100A ¼ 0:1275 100
0:0494 ¼ 15:575 mm da

From Eq. 14.33, downcomer backup, hL;dc ¼ hw þ how þ hdc;prdrop þ htray þ D ¼ 73 þ 29:2 þ 15:575 þ 104:46 þ 40 ¼ 262:235 mm  Considering check of Eq. 14.34, hL;dc  300zTS 2 (Ok) From Eq. 14.35, residence time in downcomer A h tdc ¼ dcQLL;dc ¼ 0:6372
0:262 0:0151667 ¼ 11 s (sufficient) Downcomer velocity ¼ 0:0151667 0:6372 ¼ 0:024 m s (sufficiently below choking limit) Throw of liquid over weir into downcomer is not important in case of segmental downcomers and hence omitted. Weep holes  For each tray, An;L ¼ A  Adc þAo;tray ¼ ð0:88
5:31Þ  3:25 ¼ 1:423 m2 An;L summed over all trays is 1:423
8 ¼ 11:38 m2

510

Chapter 14 Column and column internals for gas

As per the general recommendation (275e280 mm2 of weep hole area per m2 of net open liquid tray area summed over all trays), we consider total area occupied by weep holes as 277.75 m2 per m2 An;L . This gives total Awh in the tower as Awh ¼ 11:38
277:75 ¼ 3160:795 m2 And on each tray, Awh ¼ 395:1 mm2 Assuming dwh ¼ 3 mm, number of weep holes per tray Nwh ¼ Problem 14.2.

395:1 p=4
ð3Þ2

¼ 55:89z56 per tray.

Design of a valve tray for the top tray condition specified in Problem 14.1. mV ¼ 68495:42 kg=hr ¼ 19.026 kg=s mL ¼ 47963:22 kg=hr ¼ 13.323 kg=s rV ¼ 4:18283 kg=m3 rL ¼ 878:2 kg=m3 47963:22 ¼ 54:6 m3 =hr z 0.015171 m3 =s 878:2 68495:42 ¼ 16375.3774 m3 =hr ¼ 4.55 m3 =s QV ¼ 4:18283 QL ¼

Tray diameter Assuming a single-pass cross-flow tray and j ¼ 70%; TS ¼ 600 mm (2400 ); SF ¼ 1 and rV > 2.7 kg/m3, CAF from is  Eq. 14.72b   CAF ¼ SF
0:1392 þ 7:66
104
TS  3:906
107
TS2      1:56075
106
TS
rV ¼ 0:1392 þ 7:66
104
600  3:906
107
600
600   1:56075
106
600
4:183 ¼ 0:454: From Eq. 14.68 rffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi rffiffiffiffiffiffiffiffiffiffiffiffiffiffi rv 4:183 ¼ Q ¼ 0:31477 m3 =s z 0.315 m3 =s Qadj ¼ 4:55 v v 878:2  4:183 rl  rv Assuming single-pass tray and using Table 14.12, we get from Eq. 14.69, adj adj 2 1:719 ðDapprox Þ2 ¼ 9:924Qadj v þ 62:4Qv QL þ 630:205Qv QL þ 433:1QL

Dapprox ¼

ffi rffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi 





adj adj 2 1:719 þ 62:4
Q þ 630:205
Q þ 433:1
Q 9:924
Qadj
Q
Q L v v v L L

qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi ð9:924
0:315Þ þ 62:4
ð0:315
0:0151667Þ þ 630:205
0:315
ð0:0151667Þ2 þ 433:1
ð0:0151667Þ1:719 ¼ 1.947 m

From Table 14.13, the estimated Dapprox does not confirm to single-pass tray. So we repeat our calculations for two-pass tray. From Eq. 14.69 using Table 14.12,

14.7 Design illustration

ðDapprox Þ2 ¼ 9:924


Dapprox

pffiffiffi ¼ 2


511

 1:719 Qadj Qadj
QL Qadj
Q2L QL v þ 62:4
v þ 630:205
v þ 433:1
2 4 8 2

( 9:924


0:315 0:315
0:0151667 0:315
0:01516672 þ 62:4 þ 630:205
þ 433:1 2 4 8

)0:5   0:0151667 1:719
¼ 1.866 m; 2 which is close to the preferred diameter for two-pass tray and less than the minimum diameter for three-pass tray. Flow path length From Eq. 14.70 750
Dapprox 750
1:866 ¼ 699:75 mm z 700 mm ¼ 2 NP Minimum active area From Eq. 14.71 considering 70% flooding

 3:2808
Qadj þ ðQL
FPL=224:24Þ ð3:2808
0:315Þ þ ð0:0151667
700=224:24Þ v ¼ Aa;min ¼ CAF
j 0:4543521
0:7 FPL ¼

¼ 3:4 m2 Downcomer design velocity Udc;design m/s is the smallest Udc;design value obtained from Eq. 14.74aec, i.e., smaller of the values obtained from Udc;design ¼ 0:17
SF ¼ 0:17 m=s pffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi pffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi Udc;design ¼ 0:007
ðrL  rV Þ
SF ¼ 0:007
878:2  4:18283 ¼ 0:207 m=s pffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi Udc;design ¼ 2:525
104 TS
ðrL  rV Þ
SF ¼ 2:525
104 pffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi   600
ð878:2  4:18283Þ ¼ 0:183 m s rUdc;design ¼ 0:17 m s (close to the recommended value). Minimum downcomer area From Eq. 14.75, 0:0151667  ¼ 0:127 m2 Adc;min ¼ QL = Udc;design
j = 100 ¼ 0:17
0:7 . and 0:11
p
D2 4 ¼ 0:11
p4
ð1:866Þ2 ¼ 0:3 m2 rAdc < 11 % of tray area. And Adc is the smaller of (i) 2
Adc ¼ 2
0:127 ¼ 0:254 m2 (calculated from Eq 14.64) or 0:11Aa ¼ 0:11
3:4 ¼ 0:374 m2 rAdc ¼ 0:254 m2

512

Chapter 14 Column and column internals for gas

Minimum tower cross-sectional area From Eq. 14.76, (Amin m2 ) is the larger of Amin ¼ Aa;min þ 2
Adc;min ¼ 3:4 þ 2
0:254 ¼ 3:91 m2 or Amin ¼

Vload 0:315 ¼ 1:27 m2 ¼ 0:78
CAF
j 0:78
0:454
0:7 rAmin ¼ 3:91 m2

Tower diameter Recalculated tower diameter D(in m) from Eq. 14.77 rffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi pffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi 4
3:91 ¼ 2:23 m D ¼ 4
Amin =p ¼ p From Table 17.8, D ¼ 2.3 mm ¼ 2300 mm (assumption of two-pass tray is OK from Table 14.13 A ¼ ðp=4Þ
2:32 ¼ 4:15476 m2 Downcomer width and area Central downcomer Ldc;central ¼ D ¼ 2:3 m ; Adc;central ¼ 0:13A ¼ D ¼ 0:13
4:15476 ¼ 0:54 m2 Width of central downcomer wdc;central ¼ 0:54 2:3 ¼ 0.235 m ¼ 235 mm (lies within 200e275 mm as mentioned in Section 14.3.1 under downcomer area). Side downcomer Considering lw ¼ 0:5  0:6D, we assume lw ¼ 0:62D which gives lw ¼ 0:62ð2300Þ ¼ 1426 mm m =hr 54:6 ¼ 38:3 Weir loading ¼ QlwL ¼ 1:426 , much below the limit and is OK. ðm weir lengthÞ 3

From Eq. 14.15, c ¼ ðldc = DÞ ¼ 0:62

wdc;side ð%Þ ¼ 50
1  cos sin1 ðcÞ ¼ 10:77% D From Eq. 14.16, rwdc;side ¼ 247:71 mmz248 mm  A 1 c sin 2 sin1 c From Eq. 14.17 for each side downcomer, dc;side A ð%Þ ¼ ð50 =pÞ
2 sin ¼ 5:8% z 6%

Adc;side A ð%Þ ¼ 2
6 ¼ 12% per tray. 0:2493 m2 z0:25 m2 and per tray A

This gives

Adc;side ¼ Active area From Eq. 14.80,

Aa ¼ A 

X

dc;side

¼ 0:5 m2

Adc ¼ 4:15476  ð0:54 þ 0:5Þ ¼ 3:11476 m2

14.7 Design illustration

513

Aa 3:11476 ¼ 0:75ðOkÞ ¼ A 4:15476 Recalculated % flooding From Eq. 14.71, approach to flooding is recalculated,

 3:28
Qadj þ ðQL
FPL=224:24Þ v jð%Þ ¼
100 CAF
Aa ð3:28
0:315Þ þ ð0:0151667
700=224:24Þ
100 ¼ 75:7% ¼ 0:4543521
3:11476 This is Ok. Flooding Qf ;v Qf ;v ¼

Qadj 0:315 v ¼ 0:4145 m3 =s ¼ 0:76 ðj=100Þ

Flow path From Eq. 14.78, length of flow path

D  Wdc;cental  ð2
Wdc;side 2300  235  ð2
248Þ FPL ¼ ¼ 784:5 mm ¼ 0.7845 m ¼ 2 NP Width of flow path 3:11476 ¼ 3:97 m 0:7845 much less than the limiting value.

WFP ¼ ðAa Þ=FPL ¼ 3

m Liquid loading per WFP ¼ 54:6 3:97 ¼ 13:754 hrðmÞ Number of valves on a tray Since D < 3.6 m, no major beams are required and truss lines are parallel to liquid flow path. So we select valve diameter of 40 mm and valve spacing as 2dvalve ¼ 80 mm arranged in a triangular base. From Eq. 14.81a     FPL  0:216 0:7845  0:216 þ 1
NP ¼ þ 1
2 ¼ 31:921 z 32 No. of Rows ¼ 0:5
Base 0:5
0:076

From Eq. 14.81b WFP  0:8
ðNo. of major beams þ 1Þ 0:146
NP 3:97 ¼  0:8ð1Þ ¼ 12:79 z 13 0:146
2 Total number of valves ¼ 416 which lies within the maximum recommended range of 130e150 per square metre of active area. Diameter of each orifice ¼ 40 mm. Total valve area per tray Avalve ¼ 1:2566
103
416 ¼ 0:52276 m2 Valves may be arranged as discussed in mechanical aspects in text. Valves per row ¼

514

Chapter 14 Column and column internals for gas

Pressure drop Dry tray pressure drop hdry Using the alternative method, hdry is given by Eq. 14.83  2 4:18283 ¼ 98:654 mm hdry ¼ 273:4 K
UV;valve
ðrV = rL Þ ¼ 273:4
ð8:704Þ2
878:2 Using Eq. 14.82a and b, the pressure drop is slightly higher compared to the calculated value. Total tray pressure drop is calculated from Eq. 14.85 for weir height, hw ¼ 50 mm and all other  2=3 parameters specified earlier as htray ¼ hdry þ 554
ðQL =lw Þ2=3 þ 0:4hw ¼ 116:4 þ 554 0:0151667 þ 0:4ð50Þ ¼ 163:19 mm 1:426 Downcomer pressure drop. hdc;clearance is usually 40 mm for hw ¼ 50 mm Adc;clarance ¼ hdc;clarance
lw ¼ 0:04
1:426 ¼ 0:057 m2 UL;dc;clearance ¼ QL =Adc;clarance ¼

0:0151667 ¼ 0:2661 m=s 0:057

From Eq. 14.86  2 hdc;prdrop ¼ 177:7 UL;dc ¼ 177:7
ð0:2661Þ2 ¼ 12:58 mm Downcomer backup From Eq. 14.87, downcomer backup is  2=3  QL rL hL;dc ¼ hw þ 554 þ htray þ hdc;prdrop
lw rL  rV   0:0151667 2=3 878:2 ¼ 50 þ 554 þ ð163:19 þ 12:58Þ
1:426 ð878:2  4:18283Þ ¼ 50 þ 26:8 þ 176:6 ¼ 253:4 mm ¼ 0.42TS So design Ok. Leakage (weeping) check Liquid level on tray ¼ hw þ how ¼ 50 þ 10 ¼ 60 mm (approximately) UV;valve ¼ sffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi   UV;valve ðrV =rL Þ1=2 ¼ 8:7

4:18283 878:2

QV 4:55 ¼ 8:7 m=s ¼ Avalve 0:52276 ¼ 0:6 m=s > limiting vapour velocity as listed in

Table 14.18. So leakage/weeping will not occur.

14.7 Design illustration

515

Valve (Ballast) tray specification summary:Customer ID: IIT Kharagpur, India Designer: XX Date: Dec. 28, 2019 Equipment ID: 11C01 (Benzene recovery column) D, Tray/Tower diameter (m)

2.3

Downcomer dimensions

wdc , mm

lw ,m

Adc,m2

A, tower area (m2)

4.155

Side

248

1.426

0.25

Aa , active area (m2)

3.115

Centre

235

2.3

0.54

Adc, downcomer area (m2)

0.5

Off centre

-

-

-

NP, number of passes ()

2

Off side

-

-

-

FPL, flow path length (m)

0.785

S Typ. even

496

2.852

0.5

WFP, Flow path width (m)

3.97

S Typ. odd

235

2.3

0.54

lw , weir length (m)

1.426

Total

731

5.152

1.04

hw , weir height (mm)

50

Average

2.576

0.52

Antijump baffles

No

Residence time, s

4.2

Draw pans on tray

e

Adc ð% of AÞ

12

Feed to tray, kg/hr (Liq./Vap.)

e

Downcomer type

Segmental, straight

Metal/thk; Valve Deck

SS410/ 16BWG

Downcomer clearance, mm

40 in each

Metal/thk; deck

CS/10 BWG

Seal pan

No

Loadings @ Tray 1 No. of trays in section

8

TS, mm

600

SF, system foaming factor

1

Vapour

19.026

kg/s 3

m /s

4.55

rv , kg/m3 Qadj v ,

Liquid

3

m /s

4.1828 0.315

CAF

0.454

kg/s

13.323

3

m /s

0.015171 3

rl , kg/m

878.2 Continued

516

Chapter 14 Column and column internals for gas

Pressure drop

Valve type/Thk., mm  Valve nos./Avalve m2 UV;valve ðrV =rL Þ

1=2

416/0.52276

, m/s

0.6

QL =lw , m /hr per m

38.3

hw , mm

50

htray , mm liq.

163.2

DPtray , mm WC

142

3

.

Koch-Glitsch V-1/1.5

Avalve , m/s Qadj v

0.6

2

Adc;clearance , m

0.057

hdc;prdrop , mm liq.

12.6

hL;dc , mm liq.

254

Problem 14.3. Design a sieve tray for the top tray condition specified in the batch distillation problem discussed in Section 11.9.

mL ¼ 1260 kg=hr ¼ 0.35 kg=s mV ¼ 1702:8 kg=hr ¼ 0.473 kg=s rV ¼ 3:2025 kg=m3 rL ¼ 867:4 kg=m3 QL ¼

0:35 ¼ 4:035
104 m3 =s z 1.4526 m3 =hr 867:4 QV ¼ 531.54 m3 =hr ¼ 0.1477 m3 =s s ¼ 28:25 dyne=cm

Tray diameter qffiffiffiffiffiffiffiffiffiffi qffiffiffiffi 0:35 3:2025 ¼ 0:045 Assuming TS ¼ 450 mm and FLV ¼ mmVL rrV ¼ 0:473 867:4 L  0:2 From Fig. 14.16, K1 ¼ 0:008 From Eq. 14.42, assuming ðAo =AÞ  0:1, K1Corrected ¼ K1 ðs=20Þ0:2 ¼ 0:08 28:21 ¼ 0:0857 20  0:5 qffiffiffiffiffiffiffiffiffiffiffiffiffi  ðrL rV Þ ¼ 0:0857
867:43:2025 ¼ 1:408 m s From Eq. 14.44, UfV;n ¼ K1Corrected 3:2025 r V

Assuming j ¼ 85% approach to flooding, UV;n ¼ 1:408
0:85 ¼ 1:2 m=s

14.7 Design illustration

517

2 Net area, An ¼ 0:1477 1:2 ¼ 0:12308 m 2 From Table 14.8, An =A ¼ 0:88 and A ¼ 0:12308 0:88 ¼ 0:1399z0:14 m This gives tower diameter qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi D ¼ p4
0:14 ¼ 0:422 mz450 mm on rounding off using Table 17.8, A ¼ 0:159 m2 . From economic considerations, we select single-pass cross-flow tray with segmental downcomer and hw ¼ 50 mm and lw ¼ 0:76
0:45 ¼ 0:342 m The diameter of perforations can be taken as d0 ¼ 5 mm and AA0 ¼ 0:1. Area distribution over tray This gives area ratios from Table 14.8 as

Adc ¼ 0:12A ¼ 0:12
0:159 ¼ 0:019 m2 ; An ¼ 0:88A ¼ 0:159
0:88 ¼ 0:14 m2 Aa ¼ 0:76A ¼ 0:76
0:159 ¼ 0:1208 m2 and Ao ¼ 0:1Aa ¼ 0:1
0:1208 ¼ 0:01208 m2 Corrected ‘approach to flooding’   From Eq. 14.31, UV;n ¼ QV An ¼ 0.1477 0:14 ¼ 1:055 m s UfV;n ¼ 1:408 m=s j¼

UV;n UfV;n

¼ 1:055 1:408 ¼ 0:7493z75% approach to flooding (Ok)

Check for weeping 15:806
28:25 ¼ 0:103 mm From Eq. 14.57, hs ¼ 15:806s rL do ¼ 867:4
5  ðm 0:1477 VÞ ¼ From Eq. 14.49, UVo ¼ VA=r 0:01208 ¼ 12:23 m s o Table 14.10 for tray thickness/d0 ¼ 0.6, m ¼ 0.6733  0 Co ¼ m
A Aa þ c ¼ 0:7205
0:1 þ 0:6733 ¼ 0:74535 From

and

from

Eq.

From Eq. 14.48, ho ¼ 51
ðrV = rL Þ
ðUVo =Co Þ2 ¼ 51


  3:2025 12:23 2
¼ 50:7 mm 867:4 0:74535

ho þ hs ¼ 50.7 þ 0.103 ¼ 50.803 mm From Eq. 14.14, how (in mm liquid) is   0:6667  mL 2=3 0:35 how ¼ 750 ¼ 750
¼ 8:37 mm 867:4
0:342 rL lw hlo ¼ hw þ how ¼ 50 þ 8:37 ¼ 58:37 mm Since Ao =Aa ¼ 0:1, condition of weeping is given by Eq. 14.58. RHS of Eq. 14.58,    7.433
104
ðhlo Þ2 þ 0.2358
hlo þ 3.52 ¼  7.433
104
ð58:37Þ2 þ 0.2358
58:37 þ 3:52 ¼ 14:75 RHS  ho þ hs (no weeping). Using the alternate method, K2 is obtained from Eq. 14.61 as

14.50,

518

Chapter 14 Column and column internals for gas

K2 ¼

32:9ðhlo Þ0:62 1:044 þ ðhlo Þ0:62

¼

32:9ð58:37Þ0:62 1:044 þ ð58:37Þ0:62

¼ 30:354

And UV0;min minimum design vapour velocity to prevent weeping is obtained from Eq. 14.60 as UV0;min ¼

K2  0:90ð25:4  d0 Þ 30:354  0:9ð25:4  5Þ pffiffiffiffiffiffiffiffiffiffiffiffiffiffi ¼ ¼ 6:7 m=s < UV0 ðNo weepingÞ r0:5 3:2025 V

Check for entrainment From Fig. 14.17, for FLV ¼ 0:045, j ¼ 75% j ¼ 0:045 < 0:1(Ok) This gives total amount entrained e (in kg mol/hr) for liquid flow rate NL ¼ (1260/948) ¼  jNL ¼ 0:07
13:29 z1 kg mol hr. 609.4437 kg mol/hr from Eq. 14.11 as. e ¼ 1j ð10:07Þ The decreased tray efficiency Eactual due to entrainment for Emv ¼ 0:7 from Eq. 14.12 is Eactual ¼ Emv

1 0:7
100 ¼ 66:5% ¼ 0:7 1 þ eEmv =NL 1þ 13:291

Liquid gradient For lDw ¼ 0:76; q ¼ 49.4642 ; cosq ¼ 0.65 using Eq. 14.16, D 450 ð1  cos qÞ ¼ ð1  0:65Þ ¼ 78:75 mm; 2wdc ¼ 157.5 mm 2 2 From Eq. 14.63, lpath ¼ D  2wdc ¼ 450  157.5 ¼ 292.5 mm wdc ¼

w ¼ 450þ342 ¼ 396 mm From Eq. 14.65, lf ¼ Dþl 2 2   4:035
104 ¼ 1:7456
103 m s From Eq. 14.67, Uf ¼ QL ðhlo
lf Þ ¼ 0:5837
0:396  58:37
396 ¼ 45:08 mm From Eq. 14.64, Rh ¼ hlo
lf ð2hlo þlf Þ ¼ ð2
58:37Þþ396

and Re ¼

Rh U f r l ml

¼ 0:04508
1:7456
10 0:42
103

3


867:4

¼ 162:52

From Eq. 14.66 and Table 14.11 which gives c ¼ 8.2246, f ¼ expð  1:0583
lnðReÞ þ cÞ ¼ 17:06 2 f
lpath
Uf2 17:06
0:2925
ð1:7456
103 Þ
103 3 ¼ 0:0344 9:81
Rh
10 ¼ 9:81
45:08
103 4 >1, Almost irreversible reaction

Conversion x→

1.0

Endothermic Rxn

Exothermic Rxn Keq 20

Table 15.7 Design features of packed bed reactor. reactor packed with solid which is either reactant (typically in extractive metallurgy) or catalyst (in · Tubular process industries) as discussed in Chapter 14 · Design · L /d > 50 for isothermal operation and L /d > 150 for nonisothermal operation of gas through packed bed preferred as upflow may cause fluidization · Downflow also preferred for liquidesolid reactors · Downflow · Catalyst loading on catalyst support grid; bed located between BTL and TTL and head space above and below packing filled with inert (ceramic) balls fitted with catalyst collector which is a perforated can welded inside the outlet nozzle · Outlet interstage heating/cooling in order to ensure that for endothermic reactions, inlet feed temperature is not · Typically too hot to compromise selectivity and outlet temperature is hot enough to maintain reasonable catalyst activity. For p

p

exothermic reactions, run away does not occur at outlet temperature

· Due to different reaction rate in individual beds, their sizes may not be equal Table 15.8 Design parameters of fluidized bed reactor. heterogeneous reactor (analogous to CSTR) · Well-mixed may be used as reactant (in fluidized combustion process), catalyst or inert powder to promote heat transfer · Solids size 50%) of the capital cost of the reactor system and maintenance of a sterile environment is essential for any bioprocess. Reactor type and operating strategy decide product concentration, amount and impurity, degree of substrate conversion, yield, sustainability and reliability of performance. The extent of a bioreaction is quantified by productivity of cell mass or a primary product. Although continuous process exhibits higher productivity for primary products, commercial bioprocesses are mostly batch systems. Apart from the advantages mentioned in Table 15.4, batch reactors are preferred over continuous ones due to genetic stability, higher productivity of secondary products and greater ease of maintaining sterile conditions. It also offers greater reliability and process flexibility and the same reactor can be used to manufacture different products. Scaling of batch fermenter is easier and relatively straightforward as long as the same temperature, substrate and oxygen concentration can be maintained in the larger vessel. Continuous processes are

15.3 Reactor design

549

used for manufacture of single cell protein, ethanol and large volume, growth-associated products like lactic acid. Modified forms of continuous processes, e.g., CSTR with recycle, is used in largescale low product value processes like waste treatment and fuel grade ethanol production. Multistage continuous systems are used for genetically unstable cells and production of secondary metabolites and fed batch system are used to overcome substrate inhibition by periodic substrate addition. They allow flexibility of operation and are widely used in commercial plants. The designer needs to trade off between improved product quality and reactor productivity with increased cost for a more complex reactor system and increased medium utilization. The final choice depends on the economics of the specific situation. Commonly used bioreactor geometries are tanks with agitator, bubble columns where sparged gas agitates the liquid broth during upflow and loop reactors where mixing and liquid circulation are induced by the injected gas, pump or a combination of both. The flowing gas carries fluid and cells up a draft tube and disengages from the liquid at the top. The ‘degassed’ liquid flows down through the annular space surrounding the tube and is once again carried up by the flowing gas as it enters the draft tube. Several variations of the basic design are in use. The advantages and limitations of stirred tanks and bubble columns are listed in Table 15.10.

Table 15.10 Bioreactor geometry: Pros and cons. Reactor geometry Issues

Stirred tank

Bubble column

Viscosity

Up to 2000 cP (2 Pa$s)

Low viscosity Newtonian broths

Bubble coalescence problem

Not severe

Severe

Level of shear

High

Low

Advantages

High flexibility Provides high kL a for gas transfer Scale-up easier

More energy efficient (higher amount of oxygen transferred per unit of power input) Low shear environment due to absence of mechanical agitation Lower cost Lower chances of contamination

Limitations

Not suitable for shearsensitive cells Stratification of contents in multiple impeller systems

Narrow range of gas flow rate as operation limited by foaming and bubble coalescence (Bubble coalescence reduced in multistage column) Less flexible design Satisfactory mixing difficult for high viscosity broth Cells tend to accumulate on bubble surface Bubble bursting detrimental to cells

550

Chapter 15 Reactors and reactor design

Tanks of volume up to 400 m3 are used as stirred tank bioreactor with L /D values usually in the range 2e3. The L/D value 1 is for animal cell culture. Disc and turbine type impeller are typically used. Marine and paddle impeller are adopted for highly shear-sensitive cellular systems. Axial flow hydrofoil impellers are also popular as they give better Design of Stirred Tank Bioreactor performance (lower power input for same level of oxygen transfer), reduced maximum shear rates and can be used for shear-sensitive cultures like animal cells and also for viscous mycelial fermentation broth. DI ¼ 0.5D is used for axial flow hydrofoil impeller and 0.3e0.4D for Rushton impeller. The impeller speed should be sufficient to disperse bubbles uniformly throughout the tank, ensure sufficient residence time and shear large bubbles while maintaining shear-sensitive (animal) cells intact. Multiimpeller systems prevent stratification. Stirrer power is up to 5 kW/m3. Baffles of width 0.08e0.1D are used for better mixing and gas dispersion. With animal cells, baffles cause shear damage and bottom drive axial impellers slightly offset from centre are used instead of baffles. Gas is supplied to tank by sparger under pressure. The sparger may be a tube with orifices or a ring with holes. Bubble size is a function of sparger dimensions and extent of gas dispersion is decided by impeller design. Considering the requirement of sterilized condition of reaction and corrosive and/or abrasive nature of several fermentation media, fermenters are generally made of stainless steel (Type 316 for wetted parts, 304 on covers and jackets) glass. For plant and animal cell tissue culture, low C steel (type 316L) is used. Bench-scale fermenters of volume 50e500 L are glass vessels with SS cover plates. Scale-up to large reactors is limited by provision for adequate oxygen supply and efficient removal of metabolic heat. Similar oxygen supply can be maintained in the larger reactor by increasing the rate of agitation or air flow rate which ensures a higher overall mass transfer coefficient kL a. Most fermentation processes are mildly exothermic and the heat is removed either by internal coils of SS or jacketed vessels. Internal coils are preferred due to the larger surface area and more efficient heat removal but in many systems, coils rapidly foul up by microbial growth that reduces heat transfer and also adversely affects mixing. In such cases, jacketed vessels are used. Foaming is a serious concern as foam escapes from reactor and wets filter, increases pressure drop and decreases gas flow. It also provides an access for contaminating cells to enter fermenter which destroys the sterile environment and causes loss of product, time and money. Foam can be controlled either by mechanical foam breaker or addition of surface-active chemical agents. Chemical agents, although more effective, lower kL a values, reduce reactor capacity to supply oxygen or other gases and at times can inhibit cell growth. All vessels require head space for vapour disengagement and foam breaking and the working volume (culture amount) is usually less than 75% of total fermenter volume. The challenges in scale-up are mainly related to transport processes and can be summarized as follows: • Similar to chemical reactor, the relative time scale for mixing and reaction decides the degree of heterogeneity in a fermenter. When the rate-limiting step changes, the results of pilot scale cannot be used to predict performance at Scale-up large scale.

15.4 Design illustration

• •

•

•

•

551

Ensuring and maintaining homogeneous distribution of phases with minimum shear is difficult in large scale while the same can be ensured in laboratory experiments. Scaling at same L /D decreases surface to volume ratio and reduces relative contribution of surface aeration to oxygen supply and dissolved CO2 removal. This is more critical for shear-sensitive cultures where stirring and sparging is restricted. The physical conditions can never be maintained same if geometric similarity is maintained and a change in scale results in a change of physical environment of the cells that in turn alters the distribution of chemical species in the reactor. This may destroy/injure cells and change the metabolic response of culture from one scale to other. In bacterial and fungal fermentation, wall growth is important. If cells adhering to surface have altered metabolism (maybe due to mass transfer limitations), small-scale fermenters may not predict culture response in large size. Change in culture due to increased culture time.

Common scaling parameters are power to volume ratio, kL a (overall volumetric mass transfer coefficient), impeller tip speed, combination of mixing time and Reynolds number or constant substrate or product (usually DO concentration) concentration level.

Sterilization • • • • • •

Pressurized steam is used for ‘in-place’ sterilization of reactor, seals, probes, valves. All ports and valves are protected by steam sterilizable closures. The number of openings for internals and probes is kept to a minimum. Small openings are sealed with ring gaskets and flat gaskets are used for larger ones. Double mechanical seals are used for moving shafts entering fermenter. Smaller reactors may use magnetically coupled stirrers. All surfaces are electropolished (particularly for animal or plant tissue culture reactors) to avoid crevices where contaminating organisms surviving sterilization may remain trapped. All surfaces are cleaned ‘in place’ by spray using highly alkaline detergents. Diaphragm valves are preferred as they isolate the process from environment and are easier to clean. Plate heat exchangers are preferred over shell and tube exchangers for the same reason.

15.4 Design illustration Reactors can be of various types. Only a typical example is included here. Two other examples, discussed in Chapter 21, are - Design of a solid catalyzed liquid-phase reactor (Section 21.2.1) - Design of an activated sludge reactor for effluent treatment (Section 21.2.2) Problem: Design a nitration reactor for a 10 tons/day mononitrotoluene plant. The reactor is for direct nitration of commercial toluene by a mixture of nitric acid and sulphuric acid to produce a crude mixture of ortho, meta and para nitrotoluene. (Note that in texts and diagrams, Rx refers to reactor and Rxn refers to reaction.)

552

Chapter 15 Reactors and reactor design

Solution Direct nitration of toluene produces a mixture of ortho, meta and para isomers of MNT. Raw materials are of commercial grade as per the following specifications Toluene HNO3 H2SO4

99 % w/w (þ 1% inert paraffin) 60 % w/w 96 % w/w

The plant capacity is 10 tpd. Additional production of 0.6 tpd from the reactor is considered to take care of product losses in steps of product washing, separation, etc. Based on literature survey, the following conditions around the reactor are set Nitrating mixed acid concentration is w75% w/w. Maximum 5% w/w of nitric acid to limit formation of di- and tri-nitro-toluene to a reasonable extent. Side reaction(s) -

Formation of di- and tri-nitro-toluene is negligible as HNO3 at reactor inlet shall be limited to 5% w/w maximum. 1% w/w of toluene get converted to nitro-cresol. Reactions Main Rxn:

C6 H5 CH3 D

HNO3

[

C6 H4 NO2 CH3 D

H2 O

MW:

92

63

137

18

tpd:

7.118

4.875

10.6

1.393

Side Rxn:

C6 H5 CH3 D

HNO3

MW:

92

tpd:

0.073

[

C6 H3 OHCH3 NO2 D

2NO2 D

H2 O

63

153

46

18

0.15

0.121

0.073

0.029

Say, y ¼ tpd of 96% w/w H2 SO4 supplied to have 75% w/w of H2 SO4 in mixed acid. Since H2 O from HNO3 ¼ 3.384; 0:96y=ð3:384 þ 0:04y þ 1:393 þ 0:029Þ ¼ 75=25, neglecting small amount of unreacted HNO3 in acid phase. r y ¼ 5.079 tpd. In 96% H2 SO4 from storage (17.164 tpd): the amount of H2 SO4 ¼ 16.477 tpd and H2 O ¼ 0.687 tpd. External streams to Rx (tpd) Feed toluene ¼ 7.337 Reacting toluene

7.118 þ 0.073 ¼

7.191

Unreacted toluene

0.073

Inert paraffin (1%)

0.073

15.4 Design illustration

553

Feed 96% H2 SO4 ¼ 17.164 H2 SO4

16.477

H2 O

0.687

Inert paraffin (1%)

Feed 60% HNO3 ¼ 8.460 4.875 þ 0.15 ¼

Reacting HNO3

5.025

Unreacted HNO3 (1%)

0.051

H2 O

3.384

Streams leaving Rx (tpd) Organic phase

10.867

MNT

10.6

Nitro-cresol

0.121

Toluene

0.073

Inert paraffin

0.073

Vapour leaving the Rx NO2

0.073

Acid phase

22.021

H2 SO4

16.470 (74.83%)

HNO3

0.051 (0.23%)

H2 O

5.493 (24.94%)

This has w75% H2 SO4 and hence OK

Estimation of mixed acid recycle requirement to keep HNO3 in acid phase at Rx inlet below 5%. x ¼ mixed acid recycle ð0:051=22:021Þ  x þ 5:076 5 ¼ ; r x ¼ 79:586 tpd 17:16 þ 8:46 þ x 100 Recycle acid ¼ 79.586 H2 SO4

59.554 (4.83%)

HNO3

0.183 (0.23%)

H2 O

19.849 (24.94%)

554

Chapter 15 Reactors and reactor design

Fig. P15.1 shows the mass balance around the Rx. It also shows the computed composition of required streams.

Feed toluene (tpd) = 7.337 Toluene 7.264 Inert paraffin 0.073

NO2 fume (tpd) = 0.073 NO2 0.073 Organic phase ex Rx (tpd) 10.867 10.6 (97.54%, 97.03 mol%) MNT Nitro-cresol 0.121 (0.99 mol%) Toluene 0.073 (0.99 mol%) Inert paraffin 0.073 (0.99 mol%)

Feed 60% HNO3 (tpd) = 8.460 5.076 Reacting HNO3 H 2O 3.384 Feed 96% H2SO4 (tpd) = 17.164 16.477 H2SO4 H 2O 0.687

Settler

Nitrator (Reactor)

Recycle acid (tpd) = 79.586 59.554 (74.83%) H2SO4 HNO3 0.183 (0.23%) H 2O 19.849 (24.94%)

Spent acid to recovery (tpd) = 22.021 16.477 (74.83%) H2SO4 HNO3 0.051 (0.23%) H 2O 5.493 (24.94%)

FIGURE P15.1 Mass balance around reactor system.

Mixer: Mixing of the acid streams is through a jet mixer. The recycled acid from a phase separator mixes with nitric acid and sulphuric acid streams pumped from storage under flow control. Gravity flow of toluene to the nitrator is from an overhead tank, with its flow rate controlled at a desired value. Nitration reactor: The product flow out of the nitrator is over an internal weir of the nitrator vessel to a settler for phase separation. Product organic phase overflows from the settler and recycle acid is taken out as the settler bottom stream (Fig. P15.1). Cooling coils in the nitrator vessel remove heat and a propeller agitator is used to ensure intimate contact of the organic and the acid phase. The nitrator body is of cast iron, made in one piece and 75 mm thick supporting lugs cast on the outside. The bottom also has pad cast on the inside to mount studs for holding the acid distribution baffle as well as the coil supports, as required. Organic phase overflows an internal weir and enters a 100 mm NB SS pipe connected to the nitrator top. The vessel has a conical/dished bottom of w75 mm depth. The toluene distributor is formed from an SS pipe of 25 mm NB attached to vertical SS pipe which passes up through the lid of the nitrator. Another 25 mm NB pipe attached to the lid carries the nitrous fumes to an alkali scrubber. The cover plate/lid is made of SS. Two openings on the lid are provided for mounting sight glasses on nozzles. The stirrer shaft passes through the seal centrally through the lid. Individual cooling (chilled) water nozzles enter and exit the lid such that the coils get

15.4 Design illustration

555

lifted out of the vessel along with the lid. This requires an external lifting winch for lifting the lid that has lifting hooks. Mixing of recycled acid and make up acid streams. Relative enthalpy (kJ/kg)

Flow rate (tpd)

Enthalpy (kJ/d)

Inflow streams 60% HNO3

202.362

8.460

96% H2 SO4

81.41

17.164

290.75

79.586

Recycled acid (w75% H2 SO4 )

26248:9  103

Outflow stream 279.12

Mixed acid to nitrator

105.21

29366:2  103

If all inflows and outflows are at 30 C, Heat released ¼ 29366:2  103  26248:9  103 ¼ 3117:3  103 kJ/d. 3117:3  103 ¼ 44:4 C ¼ w45 C, based on estiExit temperature from mixer ¼ 30 þ 105:21  103  2:055 mated acid Cp ¼ 2.055 kJ/kg. We consider operating the Rx at a safe temperature of 45 C and include a heat exchanger to cool the recycled acid from 45 to 30 C. Heat duty of the exchanger ¼ 79:586  ð45  30Þ  2:055 ¼ 2453:2  103 kJ/d ¼ 28.39 kW. Chilled water @20 C may be used as coolant in this exchanger. The design contemplates including a chilled water recirculation unit of adequate capacity to provide chilled water at 20 C for cooling of reactor, recycle acid and any other heat removal requirement. Reactor liquid holdup estimation Kinetic data are based on literature (McKinley, C., and White, RR, AIChE Trans., 40, 143 (1944)). Rate equation R ¼ k  x  N p , lb mol MNT per hr per ft3 acid phase k ¼ rate constant, function of % w/w H2SO4 in acid phase ¼ 126 x ¼ mole fraction toluene in organic phase ¼ 0.0099 N ¼ mole% of HNO3 in acid phase ¼ 0.15 p ¼ function of % w/w H2SO4 in acid phase ¼ 0.43 rR ¼ 126  0:0099  0:150:43 ¼ 0:5517 lb mol/hr$ft3 acid phase ¼ 0:5517  137  0:4536  35:317 ¼ 1210:8 kg/hr$m3 acid phase MNT production rate ¼ 10:6  103 24 ¼ 441:67 kg/hr. Hence, acid-phase volume in Rx ¼ 441.67/1210.8 ¼ 0.365 m3 r96%H2 SO4 ¼ 1842:7 kg/m3 r75%H2 SO4 ¼ 1614:7 kg/m3, r60%HNO3 ¼ 1370 kg/m3 Considered: racid phase ¼ w1842:7 kg/m3, as it has w75% H2SO4 rorganic phase ¼ w1139:2 kg=m3 ; r

Volume of organic phase ð10:867=1139:2Þ ¼ ¼ 0:152 Volume of acid phase ð101:607=1614:7Þ

556

Chapter 15 Reactors and reactor design

Volume of organic phase in Rx ¼ 0.152  0.365 ¼ 0.055 m3 Total liquid volume in Rx ¼ 0.055 þ 0.365 ¼ 0.420 m3 Reactor volume ¼ 0.420 m3 þ volume of cooling coil and other fittings inside the Rx vessel þ vapour space volume. All coils essentially need to remain immersed in liquid for effective heat removal. Heat removal from nitrator (Q) ¼ A þ B e C, where A ¼ heat of reaction @558 BTU/lb toluene (¼ 1,297,908 J/kg toluene) B ¼ heat of dilution of H2 SO4 C ¼ sensible heat picked up by toluene from storage temperature (30 C) to 45 C (this is ignored as a conservative assumption) MNT produced ¼ ð7:264  0:073Þ  103 kg/day. ð7:264  0:073Þ  103 ¼ 108;024 W. A ¼ 127; 908  24  3600 Relative enthalpy of mixed acid (105.21 tpd) ¼ 279.12 kJ/kg. Relative enthalpy of recycled acid (101.607 tpd) ¼ 290.75 kJ/kg. B ¼ ð 105:21  279:12 þ 101:60  290:75Þ  103  103 ð24  3600Þ ¼ 2037:3 W. C ¼ neglected as a conservative measure for overestimating cooling load. Q ¼ cooling coil load ¼ 108024 þ 2037.3 ¼ w110,061 W. Considering chilled water available at 20 C and returned at 35 C, and the bulk liquid in the reactor at 45 C, LMTD ¼ (25 e 10) / ln (25 / 10) ¼ 16.37 C. :

r m water ¼

110061 ¼ 1:6085 kg=s 4:18  103  16:37

Estimating h0 The cooling coils are fed with chilled water @20 C. For a helical coil in an agitated vessel, the outside coil surface heat transfer coefficient relationship is   2    0:62   Cp m 1=3 ho D mw DI NI r  ;  ¼ 0:87  k m k m where (in consistent units), D ¼ vessel diameter ¼ 1.1 m mw ¼ dynamic viscosity of the stirred liquid at coil wall temperature ¼ 8:6  103 Pa.s m ¼ dynamic viscosity of the stirred liquid at bulk temperature ¼ 6:5  103 Pa.s r ¼ density of the stirred liquid bulk ¼ 1664 kg/m3 k ¼ thermal conductivity of the stirred liquid bulk ¼ 0:398 mW 2 :K J Cp ¼ specific heat of stirred liquid ¼ 1:88  103 kg:K DI ¼ impeller diameter ¼ 0.35 m NI ¼ impeller speed ¼ 2.5 rev./s, i.e., 150 RPM The aforementioned values are for the reaction mixture bulk temperature of 45 C r ho ¼ 1210:3

W m2 :K

15.4 Design illustration

557

Estimating hi We consider  the inside surface heat transfer equation for coils with turbulent flow di ui ri > 4000). (Re ¼ mi       Cp;i mi 0:4 hi di di ui ri 0:8  ¼ 0:023  f1 þ 3:5  ðdi = dhelix Þg  ki ki mi Where (in consistent units) di ¼ coil inside diameter ¼ 0.01,986 m mi ¼ dynamic viscosity of the coil liquid ¼ 0:89  103 Pa.s ri ¼ density of the coil liquid ¼ 1000 kg/m3 ki ¼ thermal conductivity of the coil liquid ¼ 0:588 mW 2 :K J Cp;i ¼ specific heat of the coil liquid ¼ 4:18  103 kg:K

The aforementioned values are estimates for chilled water at an average temperature of 27.5 C. To arrive at a conservative estimate of hi , the cooling coil helix diameter considered is for the outermost coil that has a wall(vessel) to wall(coil) spacing of 40 mm and the coil helix is with a pitch of 50 mm. Diameter of subsequent inner helices reduces by 100 þ tube o/d, i.e., 125.4 mm. The coils are connected in parallel flow configuration and considered having same flow rates. r hi ¼ 5631:8

W m2 :K

Estimating Uc and UD 1= Uc ¼

1 1 twall 1 1 2:77  103 þ þ þ þ ; ¼ ð5631:8  19:86=25:4Þ 1210:3 kwall ð5631:8  19:86=25:4Þ 1210:3 16:34

Uc ¼ 679.2

W m2 :K ,

1=UD ¼ 1=Uc þ Rd ¼ 1=793 þ 0:00035; and UD ¼ 548.8

W m2 :K

Estimating coil length Acoil ¼

Q 110; 061 ¼ 12:25 m2 ¼ LMTD  Ud 16:37  548:8

Acoil ¼ 153:5 m, this is the minimum requirement. p  do Coil size estimate We consider five concentric coils, each 13.5 turns in parallel connection. The outermost coil wall is 40 mm from vessel wall. The subsequent inner coils also have 40 mm gap between their external surfaces. The coil (helix) pitch is 40 mm. So the height of coil stack ¼ ð13:5 þ 0:5Þ  0:04 ¼ 0.56 m. The coils are held with four radial baffles. Based on these dimensions, the helix diameters are Lcoil ¼

dc1 ¼ 1:1  2  40  103  25:4  103 ¼ 0:9946 m; dc2 ¼ dc1  2  40  103  2  25:4  103 ¼ 0.8638 m;

558

Chapter 15 Reactors and reactor design

dc3 ¼ dc2  2  40  103  2  25:4  103 ¼ 0.7330 m; dc4 ¼ dc3  2  40  103  2  25:4  103 ¼ 0.6022 m; dc5 ¼ dc4  2  40  103  2  25:4  103 ¼ 0.4710 m; Coil details Coil no.

Helix diameter, m

Coil length per turn, m

Coil pitch, m

Coilewall or coilecoil gap

Length of coil for 13.5 turns

1

0.995

3.125

0.040

0.040

42.2

2

0.864

2.714

0.040

0.040

36.6

3

0.733

2.303

0.040

0.040

31.1

4

0.602

1.893

0.040

0.040

25.5

5

0.471

1.482

0.040

0.040

20.0

Total length of coil ¼ 155.4 m > Lcoil , r OK Immersion volume of coil ¼ ðp =4Þ  0:02542  155:4 ¼ 0:079 m3 Depth of liquid level from bottom of cylindrical portion ¼ 0.56 m Depth of bottom cone ¼ 75 mm Vapour space above the liquid level ¼ 300 mm

Volume of liquid up to coil level ¼ Volume of cylindrical part þ Volume of conical bottom e Volume of coil e Volume of immersed fixtures for stirrer and its bottom support, coil holding clamps/baffles, etc., ¼ ðp =4Þ  1:12  0:56 þ 13  ðp =4Þ  1:12  0:075  0:079  0:06 ¼ 0.417 m3 < 0.420 m3, which is the liquid hold up requirement based on reaction kinetics. Thus, construction with the above specifications ensures submergence of all the coils. The vessel has the following general arrangements and features Nitrator vessel: The 1100 mm inside diameter nitrator body is of cast iron with supporting lugs outside. The outlet weir level is at the level of the top turn of the coils over which the liquid discharges under gravity into an 80 mm diameter outlet nozzle. Vapour space of 300 mm above the overflow weir elevation is provided. The nitrator has a 100 mm diameter nozzle at the bottom, fitted with a quick opening valve for emergency draining of the vessel content in case of temperature run away. Vessel cover (lid): The lid may be a flat closure with flange. The acid inlet and the toluene inlet nozzles (each 50 mm diameter) enter through the stainless steel top cover. A 75 mm nozzle on the lid serves as nitrous fume exit. Temperature probe and a pressure-vacuum gauge are mounted on the lid. The agitator shaft passes centrally through a mechanical seal. Bottom of the agitator shaft rests on a pedestal at the bottom of the nitrator vessel. The motor for the agitator with its gearbox is mounted on the lid. The cooling coil inlets and outlets enter and leave from opposite sides. A proper backing flange is used to fix the lid to the matching flange of nitrator vessel with gasket and bolts. Lifting hooks (four nos.) on the lid are used with external winch to place and remove the lid. Depth of liquid level from bottom of cylindrical portion ¼ 0.56 m. Vapour space above the liquid level ¼ 300 mm.

Further reading

559

Depth of cone ¼ 75 mm; a hand sketch may be made by the reader in line with Fig. P15.2 that shows general arrangements of a similar nitrator. The same shall serve as a basis for developing the fabrication drawing for the equipment.

FIGURE P15.2 General arrangements of a typical nitrator with cooling coils.

Further reading 1. Carberry, J. J. (2001). Chemical and catalytic reaction engineering (edition). Dover. Original McGraw HillInc. NY 1976. 2. Levenspiel, O. Chemical reaction engineering, 3rd ed. John Wiley & Sons. Inc., New York. 3. Michael, L. Shuler, & Kargi, Fikret (2002). Bioprocess engineering: Basic concepts. Prentice-Hall of India Pvt. Ltd. 4. Towler, G., & Ray, S. (2012). Chemical engineering design: Principles, practice and economics of plant and process design. Elsevier. 5. Coulson, J. M., & Warner, F. E. (1949). A problem in chemical engineering design: The Manufacture of Mononitrotoluene. Great Britain): Institution of Chemical Engineers.

SECTION

Plant hydraulics and process vessels

V

All conventional wisdom has an element of truth in it but good design requires more than an element of truth e it requires an ensemble of correct assumptions and valid calculations eHenry Petroski

CHAPTER

16

Plant hydraulics

16.1 Introduction Almost all process industries handle fluids, which commonly, are water and aqueous solutions, hydrocarbons and organic liquids, inorganic liquids, gases and vapours. This calls for a working knowledge of the various types of pumps, compressors, pipes and pipe fittings, as well as their usage. One newly initiated to the industry will come across the use of two terms ‘pressure’ and ‘head’ interchangeably in a qualitative sense. Due to the use of different codes, one may also find a mix up of units during discussions, particularly for the pipe and pipe fitting dimensions (mm, as well as inch), as well as their pressure ratings (psi, kPa, kg/cm2). In this text, although it has been attempted to use SI units, in some cases, deviations and mix up was inevitable to have relevance with industry parlance.

16.2 Pumps Pumps are bought out items. Engineering around the pump includes piping, pipe fittings and instrumentation directly related to the process liquid it pumps. It should also include special requirements of certain applications like: • • •

flushing of the mechanical seal using an external fluid cooling of mechanical seal or gland cooling arrangements for ‘hot’ pumps (for hot fluids) for its bearing, seal/gland/pedestal/casing. Fig. 16.1 illustrates a typical process pumping system using a centrifugal pump.

P

Flushing liquid

RO P Q

Q P

FIGURE 16.1 A standalone centrifugal pump with external flushing, associated piping and instrumentation. Process Equipment and Plant Design. https://doi.org/10.1016/B978-0-12-814885-3.00016-6 Copyright © 2020 Elsevier Inc. All rights reserved.

563

564

Chapter 16 Plant hydraulics

In industry, often, pumps are provided with a spare. Fig. 16.2 illustrates a typical process pumping system using a centrifugal pump with a spare. The piping allows the flexibility to change over the service from the running pump to the spare pump with continuity of process flow controlled at the desired value by the flow control loop downstream (not shown in the diagram). Industrial installations have the following typical features: • • •

•

•

•

• •

isolation (gate) valves upstream and downstream of the pumpeessential to isolate and drain the pump casing before handing over for maintenance. suction strainerearrests particulates in the incoming liquid and prevents pump damage; during commissioning and start-up, the strainers require more frequent cleaning. pressure gauge on suction and discharge line to monitor the hydraulic health of the flow circuit, e.g., clogging of lines, incorrect line up of piping, etc. The discharge pressure gauge of a centrifugal pump showing that the pump is unable to develop pressure often points to cavitation (discussed later). Unusual fall in pump suction pressure may suggest suction line chokingeoften in the suction strainer. nonereturn valve in the centrifugal pump discharge prevents backflow through the pump when either idle or on tripping due to some mal-operation; such backflow may turn the impeller in the reverse direction and loosen it from the shaft. valve and piping arrangements for changing over from running to standby or spare pump. The reader is encouraged to decide on the step by step procedure to change over from the running pump to the spare one. drain and vent arrangements - Draining the process liquid (hazardous or otherwise) is an essential step before handing over the pump and associated piping to maintenance. Venting is required to ensure that the pump casing is filled with liquid during priming. globe valve for flow regulation on the discharge line in case of a centrifugal pumpethis is not shown in Fig. 16.2 and is a part of the flow control system downstream of the flow. valve in the suctionedischarge bypass lineethis is provided only for positive displacement pumps, where, in some cases, the capacity control is effected with a globe valve fitted in this line. A globe valve parallel to a restriction orifice may also be used to ensure a minimum bypass flow in installations of a positive displacement pump.

Sometimes a spare may have arrangements for auto-start. It is also common to find a common spare pump for two services, i.e., two pumps for different services with a common spare that can be lined up for either of the services by operating the appropriate valves. As shown in Fig. 16.2, both pumps can be kept lined up, i.e., all valves in ‘open’ position and only one pump runs. With auto-start arrangement, stopping of the running pump immediately starts the idle spare pump. Backflow from the discharge header through the pumps is prevented by the non-return valves at the pump discharge end.

16.2 Pumps

565

1 2 8 7 P 6 4 P

1

2

3

P

5 P

5

M

1: Gate valve for isolation during maintenance 2: Spectacle blind for positive isolation during maintenance 3: Strainer to prevent particles enetering the pump

M

4: Pressure gauge to indicate suction pressure 5: Reducer / expander to match line size with pump nozzle 6: Pressure gauge to indicate discharge pressure 7: Non-return valve to prevent back flow through the pump 8: Bleeder (Gate) valve

FIGURE 16.2 A typical process pumping system using a centrifugal pump with facility to change over from the running to the spare without shutting off the flow.

16.2.1 Common pump types Common pump types are centrifugal and positive displacement.

Centrifugal Pump This is the most common pump in industry as it can handle liquids up to a reasonably high viscosity (600 cSt) and slurries containing up to 20% solids. These are typically used in applications that require a low head (2e100 m). The capacity of these pumps can vary from flow rates of 0.5e2
104 m3/hr and the discharge pressure ranging from few metres to 48 MPa (approx.). The capacity is controlled by the mere throttling of the pump discharge. In its simplest form, it consists of an impeller rotating within a casing (Fig. 16.3). The casing guides the liquid from the suction to the pump centre or impeller eye. The vanes of the rotating impeller impart a radial and rotary motion to the liquid and force it to the outer periphery of the pump casing, which may be circular, volute type or diffuser type. The volute gradually increases in cross section as it wraps around the casing. This results in a gradual conversion of the fluid velocity head into static pressure head, thus minimising the hydraulic radial forces, which result due to static pressure build-up in the casing. In the diffuser arrangement, a set of stationary vanes surrounding the impeller ensures a gradual increase in flow area and less turbulence as the liquid passes through the diffuser vanes. The fluid at a higher pressure is then discharged from the discharge nozzle of the pump. The impeller consists of a number of blades/vanes mounted on a shaft that projects inside the casing. Pumps where the impeller rotates in the vertical plane are more common than vertical pumps. Liquid entry to the impeller eye may be from one side (single suction), as shown in Fig. 16.3 or from both sides in Impellers and stages case of double suction. Double suction largely balances the axial force due to change in the direction of flow at the eye. This also reduces the pressure drop in the fluid approaching the eye as the flow is split into two halves. Impellers can be open type where there is no disc encasing the blades attached to the hub. Closed impellers have discs on both sides, and partially open impellers have only one disc. Most centrifugal pumps are single-stage (one impeller). High pressure centrifugal pumps may use several

566

Chapter 16 Plant hydraulics

impellers on the same shaft developing pressure in stages. Internal channels in the casing direct the flow from the outlet of one impeller to the inlet of the next, and the final head delivered is the summation of pressure developed by each stage. The number of impellers decides the pump stages. In order to cancel the axial thrust on the shaft, the fluid entry to the eye of impellers is from one side for the first half of the stages and from the opposite side in the remaining stages. A typical multi-stage pump is illustrated in Fig. 16.4. Boiler feed pumps are usually multistage centrifugal pumps. Centrifugal pumps can also be classified based on the manner of fluid flow through it. In radial flow pumps, the liquid enters at the centre of the impeller and discharges along the impeller blades at right angles to pump shaft while in the axial flow pumps (also called propeller pumps), the discharge is in a direction parallel to the pump shaft. Mixed flow pumps combine the characteristics of radial and axial flow types. A small clearance is necessary between the impeller and the casing to allow free rotation of the impeller within the pump casing. Recirculation of pressurised liquid from the periphery of the impeller to its eye through this clearance is minimised by fitting wear ring(s) that reduces the clearance to a practical low value. Wear rings are Wear ring replaceable items. The rotating shaft enters through the casing and requires a sealing arrangement to prevent leakage of the pumping liquid through the clearance between the shaft and the casing. This would happen when the pressure inside is above atmospheric. In pumps operating under vacuum, air may get sucked in and lead to cavitation. Either a Pump seals ‘gland seal’ or a ‘mechanical seal’ is used to provide this sealing. Mechanical seals are more expensive and are used for corrosive liquids, hydrocarbons, solvents, etc., and also when the pressure is high, e.g., in boiler feedwater pumps.

Discharge

Shaft Eye of impeller

Suction

Volute Impeller

Casing

FIGURE 16.3 Components of a single-stage centrifugal pump. From US Department of Energy. “DOE Fundamentals Handbook: Mechanical Science.” US DOE (1993).

16.2 Pumps

567

FIGURE 16.4 Multi-stage centrifugal pump. From US Department of Energy. “DOE Fundamentals Handbook: Mechanical Science.” US DOE (1993).

Positive displacement pumps The positive displacement pumps can be classified as reciprocating pumps, rotary pumps and diaphragm pumps. These pumps operate in a cycle and deliver liquid in discrete volumes. The discharge is approximately constant irrespective of the head at a fixed speed as long as the capacity of the power unit or strength of pump components is not exceeded. The maximum head developed depends on drive power for negligible slippage and the mechanical strength of parts. Continuous delivery can be ensured by using several chambers with overlapping delivery periods. Typical P&ID for a reciprocating and a rotary pump is shown in Fig. 16.5. Capacity control is attained through the suction-discharge bypass line or by altering the speed. The suction-discharge line must be left open while starting the pump. In some models, a relief valve across the suction and discharge may be integrated with the pump. Additional relief valves provided on the pump discharge line is set to protect the pump and piping from overpressure in case of accidental closure of the valve or blockage in the discharge line. The relief capacity is kept the same as the pump capacity. Check valves in the discharge line are installed to prevent any backflow.

Reciprocating pumps Reciprocating pumps are used for: a) Low flow rates as in metering pumps b) High discharge pressure c) Clean liquid not containing any solid particles In a reciprocating pump, a cylindrical piston, bucket or round diaphragm moves to and fro in a chamber. Valves at inlet and discharge are usually operated by pressure difference. The pumps can be single acting or double acting with one, two or more cylinders. They are powered by steam cylinders, compressed air (usually for small pumps) or electric motors with a suitable arrangement. The reciprocating pump in Fig. 16.6A is steam driven. Reciprocating pumps with very high discharge pressure are often called plunger pumps (Fig. 16.6B), where a plunger is used in place of the piston.

568

Chapter 16 Plant hydraulics

P

P

(A)

P

P

(B) FIGURE 16.5 Typical P&ID with spare allowing changeover from the running to the spare pump without shutting the flow (A) Reciprocating pump (B) Rotary pump.

Pulsation dampener or an accumulator (not shown in Fig. 16.6A) is sometimes provided on the discharge line when (i) flow pulsation cannot be allowed when it causes piping vibration Pulsation dampening

(ii) acceleration/frictional loss by pulsation significantly increases pumping power (iii) peak pulsation pressure exceeds external relief valve set pressure

A back pressure regulator is provided if the difference between the discharge and suction pressure is below 2 kg/cm2, in order to:

Back Pressure Regulator

•

Prevent spillage by reverse pressure differential, e.g., feeding to vacuum pump or pumping from a higher-pressure drum

•

Overfeeding due to acceleration head

16.2 Pumps

569

The suction piping needs to be as short as possible and sized considering the acceleration head for ðNPSHÞA as discussed in 16.2.3. A booster (centrifugal) pump may be used upstream of a reciprocating pump to provide adequate ðNPSHÞA and allow higher suction line velocities. When suction line Suction Line Arrangement diameter cannot be increased, and the acceleration head results in lack of ðNPSHÞA , suction chambers/stand tubes of adequate dimension (equal to or twice the suction line diameter) is provided just before the suction line to reduce acceleration head.

FIGURE 16.6 Reciprocating pump (A) Piston pump, (B) Plunger pump. From McCabe, Warren Lee, Julian Cleveland Smith, and Peter Harriott. Unit operations of chemical engineering. New York: McGraw-hill, 5th Edition (International Edition), 1993.

Rotary pumps In rotary pumps, one/more rotating members (vane or screw or gear) displaces the liquid from the suction to the discharge side within a stationary housing. Based on the rotating member, these are classified as gear pumps (Fig. 16.7A), screw pumps (Fig. 16.7B) and moving vane pumps. These pumps are self-priming and can remove air from the suction line to produce a high suction lift. Screw and gear type pumps have no flow pulsation. Pulsation suppression devices are used for other types of rotary pumps. In order to reduce leakage of fluid from the discharge to the suction side, only a very small clearance is kept between the rotating members and between the rotating and stationary parts.

570

Chapter 16 Plant hydraulics

(A) Casing

Suction

Discharge

Meshing gears

(B) Discharge

Suction Relief valve

Shaft

FIGURE 16.7 Schema of (A) Gear pump, (B) Screw pump.

This necessitates pump operation at low speed and liquids to be free of particulates. The pump capacity decreases due to leakage backflow with increasing differential pressure across the pump. Since the gear pump has no valves or impellers, the frictional loss in gear pumps is much less as compared to reciprocating and centrifugal pumps. This makes gear pumps suitable for pumping viscous liquids like lubricating oil, fuel oil, etc.

Diaphragm pump This is also a positive displacement pump as the diaphragm acts as the piston, which is forced into motion by mechanical linkage, compressed air or fluid from an external pulsating source. Chances of a process fluid leakage are negligible as there is no contact between the liquid and the prime mover shaft, and hence, no seal is required. As the leakage is low, these are useful for handling toxic or expensive liquids. Although slurry transportation is mostly handled by open impeller type centrifugal pumps, a diaphragm pump is used for specific applications like feeding to a jig in the mineral processing industries. It can handle higher solid concentration as compared to other pumps but has low discharge pressure and poor efficiency. It also requires check valves in the suction and discharge nozzle. Fig. 16.8 shows a typical diaphragm pump.

16.2 Pumps

571

Discharge Air bleed valve

One way valve

Relief valve

Diaphragm

Reciprocating motion

Plunger

Hydraulic fluid

One way valve Refill valve Suction

FIGURE 16.8 Diaphragm pump.

16.2.2 Pump performance and hydraulics The rate of energy transmitted to the liquid being pumped is the product of mass flow rate ðmL Þ and the total pressure differential across the pump ðDPtotal Þ. Instead of the mass flow rate ðmL Þ the product of volumetric flow rate (Q) and density ðrÞ are used for hydraulic calculations. The sum of static Pump Head and Capacity ðDPstatic Þ and dynamic pressure differential ðDPdynamic Þ gives DPtotal . The dynamic differential pressure is DPdynamic ¼ ð1=2Þru2d  ð1=2Þru2s , where r and u are the liquid density and velocity, respectively, and subscripts s and d denote suction and discharge condition, respectively. In most pumping systems, the dynamic pressure differential is small compared to the total pressure differential across the pump. In pump calculations, the pressure is conventionally expressed as the equivalent height (head) of liquid column

ðH ¼ DPrg Þ. This ensures that the characteristic Q  H curve (discussed later) is not

affected by the density of the liquid being pumped. Eq. 16.1 gives the head H(m) developed by a pump.

 1 Pd  Ps 8Q2 1  (16.1) H¼ þ ðhd  hs Þ þ 4 rg D4s gp2 Dd where Pd and Ps are the discharge and the suction pressure required to deliver flow rate Q(m3/s), hs and hd (m) are the suction and discharge nozzle elevations, Ds and Dd (m) are the pump inlet and discharge nozzle diameters, r (kg/m3) is liquid density and g (¼9.81 m/s2) is the acceleration due to gravity. The relevant notations are depicted in Fig. 16.9. Sometimes in high capacity pumps, the size of the suction nozzle is larger than the discharge nozzle to ensure a lower suction pressure drop and a lower NPSHR.

572

Chapter 16 Plant hydraulics

ud

Dd φ

Ds φ hd us hs

FIGURE 16.9 Centrifugal pump with a difference in elevation of the suction and discharge nozzle and different diameter of suction and discharge nozzles.

If the inlet and outlet nozzles are of the same diameter and located at the same elevation, Eq. 16.1 simplifies to H¼

Pd  Ps rg

(16.2)

The pump vendor requires the maximum suction pressure for deciding its seal design pressure. A conservative estimate (on the larger size) of the maximum suction pressure is the design pressure of suction vessel plus the static head due to level difference between the pump suction and suction vessel. Hydraulic calculations are done using the normal flow rate (Qnormal ). The rated flow rate ðQR Þ at suction is QR ¼ Qnormal þ Qmargin where Qmargin is decided based on the total system around the pump in order to ensure flexibility and controllability. Typical suggested values of Qmargin are 20% for overhead reflux, 10% for feed/bottom/product and 0% for other cases. For large fluctuations in the operating conditions, Qmargin can be as high as 50% or even 100% of Qnormal . For a double suction pump, ðQR =2Þ is used as the rated flow rate at suction. The minimum flow rate ðQmin Þ is determined in each individual case from the process operation details. Power transmitted to the fluid is the Fluid Horse Power, FHP (kW) ¼ H
Q
r=368, where the developed head is H (m), the flow rate is Q (m3/hr) and the liquid density r(kg/m3). Power input to pump shaft SHP (kW) ¼ (FHP)/(hp /100) for a pump efficiency of hp %. The power output of the pump motor is the same as SHP when the pump shaft is directly coupled to the motor shaft and the coupling losses are neglected, in case, a gearbox having an efficiency of hgearbox % is used to couple the motor and the pump Power shaft.    Motor Power (kW) ¼ SHP hgearbox 100 The drive motor can be single-phase or 3-phase. The use of single-phase is recommended up to 75 kW, above which 3-phase motor is used. The motor supply voltage is selected from the suggestions in Table 16.1a and the current (I amps.) is found from the following e pffiffiffi Motor Power ðkWÞ ¼ VI cos f for single-phase motor and Motor Power ðkWÞ ¼ 3
VI cos f for 3-phase motor. The power factor (cos f) can be read from the motor label or Table 16.1b may be used at the design stage. Although Table 16.2 mentions RPM for centrifugal pumps, it actually refers to the motor RPM.

16.2 Pumps

573

Table 16.1a Typical supply voltage in India (50 Hz). 230 V ± 10% 400 V  10% 3300 6600

Table 16.1b Power factor for 3-phase industrial motors. Typical 3 phase motor power factors (cos f) HP

rpm

1 2

3 4

load

load

Full load

0e5

1800

0.72

0.82

0.84

5e20

1800

0.74

0.84

0.86

20e100

1800

0.79

0.86

0.89

100e300

1800

0.81

0.88

0.91

1 HP ¼ 0.746 kW

Table 16.2 Typical rpm values for centrifugal pump. Frequency (Hz)

2 pole

4 pole

50

2960

1480

60

3550

1770

Centrifugal pumps: In the design stage, when the estimates of flow rate (Q m3/hr) and head developed (H m) are known, the following correlation may  be used to predict  pump efficiency hp in % for 15 m  H  91m and 25 m3 hr  Q  91m3 hr Efficiency

hp ¼ 80  0:9397
H þ 0:0055
H
Q  1.5137
105
Q2
H þ 0:0058
H 2 3:0284
105
H 2
Q þ 8:3464
108
H 2
Q2

(16.3)

Eq. 16.3 is similar to the equation proposed by Carl Branan (Eqn. 2.2, P25, Branan, Carl. The process engineer’s pocket handbook. Vol. I, Gulf Publishing, 1976) that predicts the efficiency within 7% of the actual value.

574

Chapter 16 Plant hydraulics

Positive displacement and Rotary pumps: Reciprocating pump efficiency is usually around 70%e85% for most industrial applications. However, for lab-scale pumps, this may be assumed to be 50%. Rotary pump efficiency can be taken as 55% during the design phase. Pump motors: Values for motor efficiency shown in Table 16.3 may be used during the design phase.

Table 16.3 Expected motor efficiency. Range of motor power, Pmotor , kW

Expected efficiency for 3 f motor, %()

Pmotor

60

Linear - 2% leakage

50 40 30 20 Equal percentage - 2% leakage

10 0

0

10

20

30

40

50 x=>

60

70

80

90

100

FIGURE 19.14 Control valve characteristics (manufacturer’s characteristics/inherent characteristics).

be partially open to regulate the flow. These valves (linear and equal percentage) may have up to 2% of the maximum flow as leakage passing through it even if the % opening is “zero” with the stem in “shut” position. Therefore, the characteristic lines need not start from the origin of the characteristic curve and can have a positive y-intercept, as seen in Fig. 19.14 for the linear and the equal percentage cases. The general equation of the equal % valve is f ðxÞ ¼ 100  aðx=1001Þ ; when various values of a generate a family of valve characteristic curves. When such a valve has w2% leakage at “tight shut” (at x ¼ 0), a ¼ 50. Most control system analyses are based on linearity d requiring the relationship between the change in the activating signal to a control valve and the corresponding change in the flow rate being linear (proportional). This is approximately valid for small changes in the opening, even for nonlinear curves in the characteristics plot when the corresponding DP change across the valve remains small. However, when the range of change in x is substantial, there is a corresponding change in DP that results in a deviation from linearity in the relation between f ðxÞ and x. Thus, although the “manufacturer’s characteristic” or “inherent characteristic” is linear, the corresponding installed characteristic becomes “nonlinear.” A control valve with “equal % characteristics” tends to show installed linear characteristics when the DP across it increases with flow. This is due to the fact that in most installations, the pressure drop in the piping increases with flow rate and the pressure differential across the valve falls with increased % opening. Based on the above, the practical rules for selecting the control valve (manufacturer’s/inherent) characteristics are: (i) select quick opening valve whenever a flow needs to be started or shut quickly e for example, valves on automated fire water shower lines need to open quickly and valves on fuel gas supply lines to burners need to be shut quickly during an emergency.

19.6 Control valves

721

(ii) if the change in valve opening causes a maximum of 10% change in DP, opt for linear characteristics. An example of this is the control valve to regulate the flow of makeup steam for maintaining pressure in a low-pressure steam header by supplying it from a high-pressure header. This scheme is shown in Fig. 19.15. (iii) other cases should opt for valves with equal % characteristics Three-way control valves are employed for splitting a flow as in case of bypassing a heat exchanger to maintain constant outlet temperature of a stream being heated. In some cases, it may be used for mixing two flows as well. Construction of a three-way valve for flow splitting is shown in Fig. 19.16. These valves usually have much higher leakage. At extreme stem positions one of the flows is supposed to be “nil” but this rarely happens in practice, and the acceptable leakage limit is up to 10%. In reality, this does not affect the operation of the process as hardly ever there is a requirement of a tight shut of any of the flow passages during regulation. The installed rangeability of a control valve is a means to express the best working flow range limits of the control valve, say from 15% to 85%. Beyond this flow range limits, the installed valve characteristics deviate from the desired characteristics. It would be mechanically difficult to perfectly machine the valve plug and seat to give the desired characterized flow when the valve plug is just Rangeability lifting off its seat. Rangeability of a control valve may be defined as the ratio of its maximum controllable flow to its minimum controllable flow. The inherent rangeability, a property of the control valve alone, may be defined as the ratio of its maximum Cv to its minimum controllable Cv between which the valve gain (slope of its Cv versus x characteristic curve) does not deviate from a specified gain by more than some stated tolerance. The valve rangeability is reduced during actual operation due to the valve pressure drop varying with flow rate and opening %. Thus, if the inherent rangeability is stated as 50, and if the valve pressure drop is 30% of the pipeline system pressure drop (DP), the installed rangeability can be expected to be only 50  O0.3 or 27. The manufacturer can only state the inherent rangeability. Typically a control valve provides good flow regulation in opening range of 30%e75%. Beyond this range, the change in flow is less sensitive to change in opening. In other words, closing the valve below 20%e25% to reduce flowrate may not be very effective and more so in old control valves that have

High pressure steam header Supply

To HP consumers

PC

SP PT

Supply

To LP consumers Low pressure steam header

FIGURE 19.15 Pressure control in a low-pressure steam header by makeup flow from high pressure header.

722

Chapter 19 Plant instrumentation and control

FIGURE 19.16 Construction of a 3-way control valve for splitting a flow.

higher leakage. Similarly, opening a valve beyond 80% in most cases will not increase the flow rate satisfactorily. Steps in sizing of control valve: (i) List inputs • Flow rate: minimum/maximum • Flow condition: Temperature, Pressure, Maximum allowable pressure drop across the valve • Fluid properties e Gas/Liquid, Density at flowing temperature, Corrosive or not • Material, size and pressure rating of the pipe to which the valve is to be fitted. Decide the fitting type e usually flanged. (ii) Decide valve characteristics based on application requirement; Decide on the material of construction (iii) Estimate Cv required for maximum flow rate; Look through literature from control valve manufacturers, for example, Masoneilan Control Valve Sizing Handbook (PDF copy widely available on the internet) to select the available valve with same or next higher Cv. Rangeability assessment: This requires the hydraulic characteristics of the line (circuit) in which the valve is to be fitted. A plot of pressure drop (excluding the valve) versus the flow rate is generated. This plot is used to estimate the DP across the valve at a different flow rate.

19.7 Instrumentation for safety

723

Estimate % opening for the chosen valve for minimum flow, normal flow, and maximum flow to assess rangeability. (iv) Finalize the valve by fixing the valve flange size that needs to match the pipe flanges. It may be noted that often the valve flange size may come out to be one (or even two) sizes smaller than the pipe size on which the valve is to be fitted. This requires the use of “reducer” and “expander” to mount the valve. In general, these fittings are concentric but when there is a possibility of vapor coming in with liquid flow or condensate coming in a vapor line, “eccentric” fittings are used to avoid the accumulation of vapor and liquid respectively. The oblique bulge in the eccentric fitting is placed at the top and bottom for vapor and liquid service, respectively.

19.7 Instrumentation for safety Engineered safety is most often implemented by, including alarms, trips, interlocks, and over-rides in the control schemes. Alarms are provided to draw the attention of the plant personnel. Alarms may be at a single level or at multiple (usually two) levels. A column bottom liquid level may have a “low level” alarm (LL) that may trigger an annunciator to draw the attention of the operator. In addition, a second level of alarm called low-low alarm may also be present to warn Alarms against a further lowering of the level and in most cases, will be actuating a “trip” to stop the column bottom pump and prevent it from getting damaged. Obviously, the LLL and similarly HHL alarms and trips are provided when ignoring the low or high alarm can lead to more serious consequences like unsafe condition or serious damage to equipment. Trips are arrangements for automated safe actions like stopping of a pump, opening of depressurization line or a drain line, etc. The requirement of action is detected by the appropriate sensor that may detect situations like “exceeding the safe filling level in a vessel,” “overpressure in a vessel” or “temperature runway in a reactor,” etc. In Trips case the action requires a pneumatic valve to be opened or closed, the same is usually done by simply installing an electrically operated solenoid valve in the air signal line to the valve. Allowing or venting out the signal pressure through the solenoid generates the required safe action on the valve. Safe operation of some processes may require following a specific sequence of operation. An example of this is the starting of the coolant flow in a nitration reactor before charging the reactants. This can be ensured by providing an interlock between the coolant line flow valve and the control signal to the reactant flow valve. Interlocks Unless the coolant flow valve is “open” the interlock arrangement will not allow the “reactant” inflow valve to be opened. There can be interlocks that may involve several parameters. Such interlocks can be implemented by electrical/ pneumatic relays, as well as PLCs and DCS. Overrides are easily understood by considering an example. The catalyst line connecting the reactor and the regenerator in a fluidized bed catalytic cracker unit (FCCU) has a control valve whose opening is manipulated to control the level in the fluidized bed reactor. The pressure drop across this valve is not allowed to fall below a Over-rides limit of w0.3 kg/cm2 to avoid flow reversal through it. In case the pressure

724

Chapter 19 Plant instrumentation and control

differential falls below the safe limit, the valve is to be closed. This is handled by signals from the level control loop being compared with a differential pressure control loop and the lower of the two selected through a low signal selector (LSS) is sent to the control valve. Thus, the signal from the differential pressure control loop overrides the level control loop signal. Similar override control is also used in centrifugal compressors that require a minimum flow through the machine to avoid surging.

19.8 Distributed control system (DCS) With the advent of microprocessors and the integration of computers in control systems, the concept of employing several individual single loop controllers for large plants have given way to DCS. There are formal definitions of instrumentation structure calling it Level - I, II, III, etc., but in essence this refers to conceptually layered structure with (i) field inDCS struments in the lowest level, (ii) controllers in the level above and (iii) a level of supervisory functions like optimization and MIS report generation in the top level. Field instruments essentially measure and transmit the process variables, and the final control elements manipulate the flows. The controllers receive their setpoints from the supervisory level either through the operator settings from the control panels, which are also called operator stations or from the optimization algorithm that may run in the supervisory level. The optimization and the MIS reporting capability of DCS make the plant operation more economical, safe, and better monitored. A typical hierarchical structure of DCS is illustrated in Fig. 19.17. This shows the Engineering Workstation (EWS) for configuring and programming the system, Operator Console(s) for operator-interface, History Module(s) for plant parameter and event history data. The connected printer(s) print the alarms, as well as the process parameter log. Typical vendors for the DCS systems are M/s Honeywell Inc., M/s Yokogawa, M/s AlleneBradley, etc. Engineering workstation

Operator Console(s)

History module

Suprvisory Computer

Printer

DCS Hardware

SP SP PV Panel mounted PV Controllers OP OP

Field instruments

FIGURE 19.17 DCS structure.

19.9 Control schemes for common processes

725

Since the DCS offers high computational power, in modern plants apart from the single loop control algorithms several options of multivariable control algorithms, are also implemented in large plants. These algorithms take care of the control objective in the presence of constraints on process parameters. Specific vendors offer these control schemes and implement on various DCS platforms. However, these are not within the scope of this book. Programmable Logic Controllers (PLCs) were developed as a flexible version for a set of relays and timers housed in a single body of electronics that can be configured. The requirement came from processes that required event-driven actions and actions that were to be taken upon certain scheduled and process conditions getting fulfilled. This PLC is in contrast to the continued steady-state desired for many plants that require very little change in operation under normal circumstances and the normal choice in most such cases fall on a DCS. PLCs are, therefore, more common for processes that require changing its process conditions/actions like charging of raw material, discharging of reactor hold up after the processing and similar situations. However, modern PLCs have developed substantially and are quite capable of operations that are also done on DCS, for example, running of control algorithms. Large plants like refineries today use DCS and the PLCs together with most discrete-time operations being carried out with the help of the PLCs and the rest being carried out by DCS.

19.9 Control schemes for common processes 19.9.1 Distillation control and instrumentation This involves control of T, P, L, F parameters in order to indirectly control the composition of the streams leaving the distillation plants. The process consists of - distillation column e with or without side stream draw - overhead condenser and reflux drum - reboiler A typical P&ID is shown in Fig. 19.18. This is for a distillation column with its auxiliaries for separating a mixture of hexane and benzene. It may be noted that the purpose of the diagram is primarily to represent the control schemes, the piping details, particularly strainers, blinds, etc., are not included. Table 19.4 lists the instrumentation legends in typical P&IDs. The distillation column controls are: • • • • •

TI for the top, bottom, and other trays PI shown for the column bottom (it may be in several other sections in the column, including on draw off trays) LIC for column bottom level manipulating the bottom product (Benzene) flow rate to storage TRC for the column top temperature controls the same by manipulating the setpoint of the reflux flow controller setpoint. This is a master-slave cascade control configuration. TRC for the third tray from the bottom controls by manipulating the steam flow to the bottom reboiler. Often the temperature of a tray close to the bottom, as in this case, is controlled instead of column bottom temperature. This is opted for columns, where such temperature is relatively more

726

Chapter 19 Plant instrumentation and control

Fuel Gas

TRC

FAL FAH FRC

FY

TT

FT

FCV

Nozzles / connections

7

PT

7

PAL

TI

PRC

TE

1. CW in 2. CW out 3. Medium pressure steam 4. Condensate drain 5. Sampling port 6. Drain to closed system 7. Vent to flare 8. Drain to plant open sewer 9. Vent to atmosphere

9

PAH

FE FI

PY

IOCO1

TI

PCV

TI

6

5

7

TRC

FEED

TI

TI

IOEO1

2

TAH

1

PI

TI TAL

5

5

LAH LE

5

3

TE

PI

LAL

IOVO1

5 TT

TI

1 PI

LIC

LT

LY

FI

6

LCV

HEXANE

FI TY

7 LE

TI

TCV

IOPO1B

IOPO1A

9

PI

PI

3 LT

IOEO2

6

6

6

4 PI

8

FI

TI

LIC

BENZENE

IOPO2A 6

TI

IOPO2B

LCV LY

6

FIGURE 19.18 P&ID of a distillation column with reboiler, condenser, and reflux drum.

• •

sensitive to changes in column bottom product composition. By controlling this temperature, the bottom product composition is effectively controlled. Column pressure is controlled indirectly by controlling the reflux drum pressure controller that manipulates the noncondensable (Fuel Gas) venting from the same. Liquid level in the overhead reflux drum is controlled by manipulating the top product (Hexane) flow to storage through the LCV

There are high and low alarms (AH, AL) provided to warn the operator in advance against abnormal situations: • •

•

Column bottom leveleLIAH, LIAL. Column bottom level touching the bottom tray can dislodge the tray. Loss of column bottom level can damage the bottom product pump. Reflux flow rate e FAH, FAL. The excess flow of reflux can signify an underperforming condenser. High flow can also flood the top tray. The low flow of reflux can lead to inadequate fractionation in the rectification section of the column and a column upset. Reflux drum pressureeThis indirectly controls the column pressure. PAH, PAL: Uncontrolled high pressure leads to high make of Fuel Gas at the cost of top product d this can happen due to inadequate performance of the condenser. Low reflux drum pressure, particularly its falling below atmospheric pressure can be potentially unsafe as it would draw air through any leakage in the

19.9 Control schemes for common processes

727

Table 19.4 Legend for instruments in P&ID. Symbol/designation

Item

FAH

Flow alarm high

FAL

Flow alarm low

FCV

Control valve manipulated by the flow controller

FE

Flow (sensing) element

FI

Flow indicator

FRC

Flow recording controller

FT

Flow transmitter

FY

Transducer to convert flow controller output signal to pneumatic signal for the control valve

LAH

Level alarm high

LAL

Level alarm low

LCV

Control valve manipulated by level controller

LE

Level (sensing) element

LI

Level indicator

LIC

Level indicating controller

LT

Level transmitter

LY

Transducer to convert level controller output signal to pneumatic signal for the control valve

PAH

Pressure alarm high

PAL

Pressure alarm low

PCV

Control valve manipulated by pressure controller

PE

Pressure (sensing) element

PI

Pressure indicator

PRC

Pressure recording controller

PT

Pressure transmitter

TI

Temperature indicator

TAH

Temperature alarm high

TAL

Temperature alarm low

TCV

Control valve manipulated by temperature controller

TE

Temperature (sensing) element

TI

Temperature indicator

TRC

Temperature recording controller

TT

Temperature transmitter

TY

Transducer to convert temperature controller output signal to pneumatic signal for the control valve

728

•

Chapter 19 Plant instrumentation and control

system. Such leakage in real plants through valve glands, pump glands, flanged joints cannot be totally avoided. Accumulation of air in the system with hydrocarbon leads to an explosive mixture. Reflux drum level e LAH, LAL. Overfilling of reflux drum may send liquid in the Fuel Gas line. Also low level in the drum leads to loss of suction of the pumps taking suction from the drum and its consequential damage. Other points to note in Fig. 19.18 are

•

•

Installation of control valves with two (gate) valves upstream and downstream and a bypass (globe) valve. This allows isolation of the control valve for calibration, maintenance, and replacement without affecting the process. During isolation, the flow is regulated manually by the globe valve in the bypass line. In many cases, blind flanges are additionally provided to guard against any accidental opening or passing of the isolation gate valves. Installation of pumps in paralleleThe arrangement allows one of the pumps to be run while the other remains stand by. This also allows changeover from one pump to the other while the plant is in operation. The nonreturn valve on the individual pump discharge line prevents backflow through the pump during the changeover. Individual PIs show the pressure developed by each pump and verifies its health.

19.9.2 CSTR instrumentation and control Fig. 19.19 shows the control and instrumentation for a CSTR fed with two liquid reactants F1 and F2. The reaction produces a liquid product (P) pumped out of the system. The product is drawn by the bottom pump, and a part of it is cooled by cooling water in a heat exchanger and recycled back to the

SP SP M FC

TC TT

FT F1 SP

RC

SP

CW SP

FC FT

LT LC

F2

P

FIGURE 19.19 Instrumentation and control system of a CSTR.

Further reading

729

reactor to remove the heat generated in the reaction. The reactor temperature is controlled by TC that manipulates the opening of the control valve for the recycle flow through the cooling heat exchanger. The flow rate of Feed F1 can be independently set, while the ratio controller (RC) sets the setpoint for the flow controller for F2. Liquid level in the reactor is controlled by the controller (LC) manipulating the opening of the control valve regulating the discharge to storage. The P&ID shows pneumatic control valves and controllers (note the symbols).

Further reading Venczel, K., & Lipta´k, B. G. (1982). Instrument engineers’ handbook: Process measurement. Chilton Book Company. Spink, L. K. (1978). Principles and practices of flow meter engineering. Foxboro, Massachusetts: The Foxboro Company.

CHAPTER

Engineered safety

20

20.1 Introduction A designer’s mandate implicitly includes “safety” as an essential component. This is in addition to the requirement of discharging the desired function of the process plant being designed and built. The designed plant must be safe to construct, operate, maintain, and also to dismantle at the end of its useful life. Safe practices and mechanical integrity are essentially built into the codes and standards. Strict adherence to applicable codes and practices not only standardizes the equipment but also ensures the mechanical integrity of the equipment and implementation of safe practices. During operation, adopting “Standard Operating Procedures (SOP)” minimizes the chances of manual error and makes the plant operation safer and efficient. The term “inherently safe” is often misleading. The inherently safe system in its literal sense of “absolutely safe” is, however, a utopia, as a small risk will always remain, however small it is. The term “intrinsically safe” used by electrical engineers refers to electrical systems with a negligible chance of generating a hazard and should not be confused with “inherently safe.” During the designing of the process equipment, plant or system facility, features to improve safety are incorporated. This can be through the incorporation of alarms, trips, override switches, limit switches, relief under overpressure or excess vacuum or by incorporating process features like overflow, inert gas blanketing and purging, positive pressurization to avoid sucking in of ambient air, cooling sprays and a host of other options. Safe disposal of hazardous discharges through the appropriately designed flare and blowdown systems is another aspect that is taken care of during the plant designing stage. Features like “safety insulation” or even guards protecting against hazardous moving machinery components and parts for personnel protection are also taken care of. All of these collectively come under “engineered safety,” and its scope and purview do not end at the design stage. Postdesign, the facility is reviewed prior to commissioning and start-up and also at every stage of any change in process, operation procedure, or facility. It is this “engineered safety” we are concerned about in this chapter. Safety in general concerns protecting people, plants, and equipment, as well as the environment, with the best approach being to detect potential unsafe conditions and guard against those. Failure to protect can have grave financial and societal consequences apart from severe operational restrictions, production loss, etc. Plants that are designed “safer” attract low insurance charges. In India, the document that covers this aspect is the “Tariff Advisory Committee Report” published by the Ministry of Finance that includes even technical “safe recommendations” and its effect on the insurance tariff. Process Equipment and Plant Design. https://doi.org/10.1016/B978-0-12-814885-3.00020-8 Copyright © 2020 Elsevier Inc. All rights reserved.

731

732

Chapter 20 Engineered safety

20.2 Hazardous area classification Process plants often handle chemicals that are hazardous. Volatile and flammable chemicals pose explosion and fire hazards. Electrical fittings and equipment are inevitable in plants, and these have the potential to be a source of ignition or explosion. Hazardous areas are those where an explosive gas atmosphere is expected to be present in quantities, such as to warrant special precautions for the construction, installation, and use of electrical apparatus. The hazardous areas have classification of zones to minimize this threat. In Hazardous Zones India, the classification of Zones is defined in the BIS standard- Classification of hazardous areas (other than mines) having flammable gases and vapors for electrical installation. The zone boundaries are defined in the code with respect to the source and considerations of the environment around the same. This code is applicable for process plants where there may be a risk due to the presence of flammable gas or vapor, mixed with air under normal atmospheric conditions. This, however, excludes mining, explosive processing, and manufacturing, areas with ignitable dust/fibers hazard, areas where pyrophores (i.e., substances that may ignite on contact with air or water) may be present, areas where ignition sources apart from electrical may be present, medical rooms, and domestic premises. All electrical equipment to be used in these zones must have appropriate design and certification. The process plant designer knows the best about the plant layout, process, and chemicals involved. It is important to arrive at safer plant layouts with a bare minimum area of the plant falling in the highrisk zone. This calls for familiarity with the said BIS code, IS 5572(2009), that also has equivalent codes practiced in different countries. Hazardous areas are classified in zones based upon the frequency of appearance and the duration of an explosive gas atmosphere as follows: Zone 0: An area/place in which an explosive atmosphere is present continuously or for long periods or frequently. Zone 1: An area in which an explosive atmosphere is likely to occur in normal operation occasionally. Zone 2: An area in which an explosive atmosphere is not likely to occur in normal operation but, if it occurs, will persist for a short period only. Normal operation of a plant refers to a situation with the operating parameters remaining within design limits at all times, including start-up and shutdown operations. The minor release of flammable material may be part of normal operation, for example, a small amount of pumping fluid (inflammable) continuously passing out of pump seals to provide seal lubrication. Failures like spillages caused by accidents that require urgent repair or shutdown are not considered as part of normal operation, nor are they considered to be catastrophic. Typically a Zone 1 area will be surrounded by a larger Zone 2 area. The possibility of any larger but infrequent release would require a larger Zone 2 area. Explosion hazard: The Lower Explosive Limit (LEL) and the Higher Explosive Limit (HEL) of an inflammable vapor refer to the minimum and the maximum concentration of the species mixed with air at atmospheric pressure within which the explosion reaction (uncontrolled Explosive limits for vapor-air mixtures combustion) would propagate in the presence of an ignition

20.3 Trips and alarms

733

source. One must note that these are indicative values and the actual limits depend on the confinement, ventilation, combination of chemical species present, etc. HEL is also known as Upper Explosive Limit (UEL). Flammability of a liquid is related to its Flash Point, which is an experimentally measured property. Lower flash point indicates a higher tendency of the liquid to form a flammable mixture with air. Flammable liquids have a flash point below 93
C and a vapor Liquid classification based on Flash Point pressure not exceeding 2.81 kg/cm2 at 37.8
C. There are three classes of flammable liquids. Class A: Flammable liquids having flash point below 23
C. Class B: Flammable liquids having flash point 23
C and above but below 65
C. Class C: Flammable liquids having flash point 65
C and above but below 93
C. The density of the saturated vapors of these flammable liquids at ambient temperature is generally less than 1.5 times that of air. In most cases, under ‘open air’ conditions a direct relationship exists between the grade of release and the zone type to which it is classified: Continuous grade normally leads to Zone 0; Primary grade normally leads to Zone 1 and Secondary grade normally leads to Zone 2. However, one may note that situations like poor ventilation may lead to a more stringent zone while the converse is true with provision for high ventilation. Sources may also have a dual grade of release with a small continuous or primary grade and a larger secondary grade. In the case of gas mixtures, the LEL or UEL can be estimated as follows e Let P1, P2, P3, etc., represent the percentage volume of each combustible gas in a mixture ignoring air and inert gas(es) present: P1 þ P2 þ P3 þ . ¼ 100; N1, N2, N3, etc., represent (in % v/v) either the lower or the upper flammable limits in the air for the individual combustible components. The flammable limit (L % v/v) of the mixture is given by - L ¼ 100 / (P1 / N1 þ P2 / N2 þ P3 / N3 þ .)

20.3 Trips and alarms Alarms draw the attention of the operating personnel to an event that triggered it. It can be visual and/ or audible annunciation. It may communicate a normal operational event like emptying of a vessel, or it can demand emergency actions to be executed manually or by an automated process. Trips associated with alarm start a procedure. Such an action can be automating start/stopping a pump in maintaining its normal and safe operation. Alarms are associated with the upper or lower limiting value of a process parameter that can be pressure, temperature, level, or flow. An alarm can also be set on the rate of change of the stated parameters or even combinations of the same. Inbuilt facilities in the Distributed Control Systems (DCS) and Programmable Logic Controllers (PLC) allow the configuring (setting) of such complex alarms and trips. Hard and Soft AlarmeAlarms and trips that are supported by individual/dedicated hardware are “hard. ” “Soft” alarms are configured through the software on DCS or PLC. Soft alarms and trips are susceptible to system-wide failures, such as failure of computers or power failure. In most cases, the hard alarms and Hard and Soft Alarms trips provide redundancy over their “soft” counterparts and is a method of

734

Chapter 20 Engineered safety

improving reliability for critical applications. This is more common for trips that are warranted for additional safety. A tilting mercury switch is often used as a dependable sensor for the level limit used for actuating hard level alarms and trips. The most common form of alarm is triggered by exceeding the preset (high/low) limit of a parameter that denotes an undesirable process condition. Its output actuates relays that may activate annunciator, light indicators on alarm panel, trigger shutdown actions, etc. Normally, a deadband is associated with the preset limit and the alarm “gets reset” once the parameter returns to its normal state. A relay contact can be normally open (NO) or normally closed (NC), and this is usually chosen such that “power failure” leads to a “safe” state. High, High-High, Low and Low-Low alarm settings: Often, two stages of alarm conditions are warranted. The first stage referred to, as the High alarm setting and the Low alarm setting are associated with less severe condition while High-High and Low-Low settings are installed as alert for more serious consequences, which Alarm settings often warrant automated tripping of some machines or actuating automated shutdown. For example, when the bottom level of a distillation column falls below say 30%, the Low-level alarm may be triggered. The Low-Low alarm for this level may be set at 10%. This Low-Low alarm also trips the bottom product pump to protect it from losing suction/ running dry. Interlocks are configured to make operations mutually dependant. An example of this may be a pump that empties a process vessel. The starting switch of the pump can be interlocked to the level sensing in the vessel. This would allow starting the pump only if the level sensor instrument in the vessel sends a signal that the level is above a minimum value. In a vessel containing hot hydrocarbon liquid under high pressure, opening the drain valve is safe after its depressurization, as this avoids the potentially unsafe condition of flammable vapor release from the flashing of the draining liquid. To ensure this, the opening of the drain valve may be interlocked with vessel pressure sensing that would avoid the accidental opening of the drain without properly depressurizing the vessel. Interlocks are easily configured on DCS or PLC using their configuration software. Dedicated hardware may be used for high reliability.

20.4 Blowdown and flare Process plants generate the reject streams that can be gas, liquid, or a mixture of both. These streams are sometimes benign but can often be toxic, inflammable or hazardous due to extreme temperature and pressure or some other property like solid content or high BOD/COD. Appropriate arrangement for safe disposal of these streams is planned during the plant design.

20.4.1 Blowdown The suitable disposal method for a discharge stream depends on its hazardous nature, frequency of discharge, discharge rate and duration, and plant economy. There may be continuous, intermittent, and/or emergency discharge streams. A continuous discharge may be bleeding of the process liquid in small quantity from pump glands/ seals, venting of compressor seals, blowdown from cooling tower basin to limit solid concentration in

20.4 Blowdown and flare

735

the cooling water, small streams of cooling water to cool pedestal/bearings/seals in pumps for hot service. The disposal arrangements can be diverse and often depend on the quality and flow rate. The cooling tower blowdown is routed to an effluent treatment plant. Inflammable vapor flow from compressor seal venting is discharged to the atmosphere through a high point vent (usually higher than any nearby shed/building), through a flame arrester. Intermittent streams are generated from time to time during routine procedures, for example, draining of pump casing before handing over for maintenance or its venting in case of vapor locking, depressurization/draining of vessels, and other equipment, etc. Emergency discharges are streams generated during abnormal operations (plant emergency situations), such as popping of safety valve, bursting of the rupture disc, accidental overflow from vessels, etc. Vapor blowdown can be to (a) atmosphere, (b) lower pressure process vessel, (c) closed pressure relief system and flare header. (a) Atmospheric venting is allowed if the pollution standard and the noise level limits can be respected, for example, low volume atmospheric discharge of noninflammable, nontoxic, noncorrosive vapors lighter than air, such as steam or nitrogen. Even H2 and light hydrocarbons (lighter than air) can be discharged in small quantities when there is no nearby source of ignition. Atmospheric discharge is not permitted if there is any risk of condensation or settling down of flammable/corrosive/toxic material that may reach the ground/lower elevation. For safe disposal, the discharge streams are often diluted with inert gas, usually steam, upstream of the discharge point. An example of this is the steam dilution of the seal vent stream from the hydrocarbon centrifugal compressor. Such discharge points are typically located at a minimum elevation not lower than the nearest plant shed or building roof. (b) Discharge into a lower pressure vesselePopping of a safety valve mounted on the higher pressure vessel often discharges vapor to another vessel at a lower pressure. This is frequently seen in cases of two-stage flashing where the safety valve of the first drum discharges into the second (Fig. 20.1A). Another example is the desalter vessel safety valve discharging into the flash zone of the main crude distillation column (Fig. 20.1B). One may note that the safety valves are mounted on vertical lines, and the downstream is kept free draining. The possibility of recovery of the discharged stream from the lower pressure process vessel is an advantage in this case. (c) Discharge to a closed pressure relief system and flareeThis is used when other methods cannot be adopted due to safety and pollution norms. The toxic or inflammable vapor released from the depressurizing lines is discharged in this way. The blowdown can be to a knockout drum from where the liquid is periodically pumped out and the vapor is safely discharged to a flare header or (rarely) vented through a high point vent. The discharge of hot vapor to the flare is often through a condenser to cool and condense part of the discharge and recover it in a knock out drum. This reduces the flare header flow, and the vapor joins the flare header at a lower temperature. In this case, the problem of thermal expansion is reduced, and the mechanical design of the flare header is simpler.

736

Chapter 20 Engineered safety

(A)

(B)

HP Flash drum Column LP Flash drum

Desalter vessel

FIGURE 20.1 (A) Safety valve of first flash drum discharging into the second drum at a lower pressure, (B) Desalter safety discharging into the flash zone of column.

Liquid blowdown streams can be discharged to (a) blowdown drum with its vapor outlet connected to a vapor blowdown system. Typical liquid residence time considered in sizing the drum is a minimum of 20 min, (b) lower pressure vessel that is required to handle the liquid, as well as the vapor generated due to flashing, In petroleum refineries, crude oil from the popping of safety valves on the crude oil electric desalter is discharged to the flash zone of the crude distillation column, (c) open sewer, in case the discharge is nontoxic, nonflammable, and safe for sewer discharge, (d) pump suction, for example, the safety valve mounted at the discharge of a reciprocating pump may discharge the liquid to its suction. In some plants e.g. the visbreaking unit of refineries, there is a requirement to handle a fairly large volume of emergency hot hydrocarbon discharge of the furnace tubes that are drained quickly to avoid coking up the tubes. This calls for a separate hot blowdown vessel with an appropriate arrangement that often includes a dedicated stack for safely disposing the vapor diluted with steam. Typically in a process plant, there are two systems to collect and handle liquid blowdownethe open and the closed blowdown arrangement. Open sewer is for infrequent, small volume liquid discharges that do not generally contain hazardous components. Plant washings containing contaminants from the process, casing draining, and seal leaks from pumps collected in pump bays, sampling port discharges for dilute aqueous solutions like acids, alkali or salts that are not too hazardous, etc., are usually drained to the sewer manifold through funnels. The open drains in the plant area have their top covered with iron grill or plates. The liquid in the drain flow by gravity and join the plant-wide sewer network and finally reach the effluent treatment plant where they are treated to remove small amounts of polluting components like oil, solvent, suspended solids, etc. Catch pits are installed at several places in the underground section of the sewer to trap any settling solid and/or any floating immiscible liquid, such as hydrocarbon oil/ solvents. In refineries, this open blowdown system is also called “oily water sewer (OWS)” as it primarily carries water mixed with a small quantity of oil. The “stormwater drainage system” is different from OWS and is primarily meant for carrying the rainwater, requiring very little or no treatment before discharge. Normally, the continuous blowdown of cooling water from the cooling tower system is routed to the stormwater drainage.

20.4 Blowdown and flare

737

Closed blowdown (CBD): Any hazardous liquid drained within the process plant area that is not safe to be discharged into open drains is handled in this system. The liquid discharged should not have high vapor pressure and should not be at high temperature to avoid a large amount of vapor generation that can pressurize the CBD system. The system consists of a direct piping connection for discharging the liquid with isolation valve(s). Opening the valve drains the liquid to the underground CBD piping that finally accumulates in the underground CBD vessel. The blowdown vessel is vented to atmosphere with steam/inert gas purge in the discharge line close to the tip, or any vapor is piped to the vapor blowdown system. The accumulated liquid is periodically emptied, usually to a tank in the effluent treatment plant by a submerged pump. Typical discharges to CBD can be liquid from sampling points, drains of vessels, or other equipment, etc. In some cases, the hot samples are cooled by cooling water in small coiled tube coolers. The design of the CBD vessel is basically sizing its volume that is based on a rational estimate of the discharge rates into it. Some inaccuracies in the volume estimate can be tolerated as in the case of a marginally small volume, the emptying pump needs to be run more frequently or for a longer time. The emptying pump capacity should be such that in normal operation, running the same for about 1 hr in a shift is sufficient. Disposal of low flow rate gas/liquid streams: a) Vapor/Gas can be discharged into the environment after adequate dilution. Discharging the seal vent stream from a centrifugal compressor after dilution with steam is common. The vent discharges typically at a minimum elevation not lower than the nearest plant shed or building roof. Another example is the gas from the sour water stripper. Stripping stream is used to strip out the dissolved H2S in “sour water” from an electric “desalter” unit for petroleum crude oil. The small flow rate of the vapor exiting the tower top containing some H2S and a little hydrocarbon may be discharged through a high point vent after dilution with steam. b) Low flow of liquid streams include casing drains from pumps collected from the pump bed, etc. Such discharges are infrequent, small in volume and may include some process liquid that although in small amount, may be hazardous. In a refinery or petrochemical plant, these are routed to “oily water sewer (OWS).” Often the sampling ports on process equipment during flushing requires disposal of a small quantity that is also sent to OWS through a collecting funnel and associated piping. If such discharges pose a threat due to high temperature, flammability, or toxicity, then instead of discharging to OWS, these are sent to a “closed blowdown (CBD)” system via a cooling exchanger, if necessary. Discharge of gas/vapor/liquid under pressure: These streams are generated mostly to relieve overpressure by the opening of the vent that discharges vapor or by popping of the safety valve. It can also be through a bursting disc that purposely fails under overpressure. Discharge of a benign vapor/ gas like compressed air or inert gas may directly be to the atmosphere at a safe location. In the case of inflammable vapor and liquid, the discharge needs to be contained, and the safety valves discharge to the flare header within the plant. The pressurized liquid flashes on discharge, and a mixture of vapor and liquid enters the flare header. The flare header piping is, therefore, sloped towards the direction of flow and all such discharges reach a knock out drum located usually at the plant boundary. The liquid from the drum is periodically emptied by a transfer pump to the effluent treatment plant. The vapor from the drum joins the main flare header and is disposed off in the flare system. Highly toxic vapor discharges are treated with an appropriate scrubbing liquid in an absorber. The scrubbing liquid, often an alkali or an acid, reacts with the toxic chemicals to neutralize them or render them nontoxic such

738

Chapter 20 Engineered safety

that the scrubber discharge gas is fit for ambient discharge. An example of this is the use of caustic soda scrubber to scrub methyl isocyanate (MIC) gas. Such a scrubber had unfortunately failed to function in the Union Carbide Plant in India, causing the Bhopal Gas Tragedy. In order to safeguard against overpressure, the designer has to choose between a bursting disc (also called rupture disc) and a safety valve. In either case, the relieving is based on the differential pressure across it exceeding the set limit, and hence, the designer looks for a nearly constant pressure location like the flare header or ambient to which the discharge is to be sent. Safety valves are supposed to relieve the overpressure by popping open at the “set pressure” and “reset” (i.e., shut) when the pressure is just below the set pressure. However, in practice, resetting happens when the pressure has fallen below the popping pressure set for the valve by a perceptible limit. In the case of “bursting disc,” the discharge continues as long as there is a pressure differential across it and the disc needs to be replaced. Normally, a bursting disc is used for relieving large gas/vapor flow.

20.4.2 Safety and pressure relief valves “Safety Valve” and “Relief Valve” are two terms often used interchangeably without considering the little differences they have. Both are for protecting against mechanical failure due to pressure build-up in the equipment/piping on which it is fitted. While the Safety valve is supposed to pop open, the pressure relief Overpressure relief valve relieves the overpressure by allowing discharge at a rate proportional to the quantity by which vessel pressure is above the set pressure. This differential quantity is also called ‘overpressure”. Different types of service (steam, air, gas, etc.) require different valve types. The material of construction also needs to be appropriate for the service. For example, valves made of stainless steel are preferable for corrosive fluids. The relieving capacity of a valve for a particular service depends on several factors, including the geometry of the valve, the temperature of the media and its orifice area. Following are the considerations for sizing Safety Valves and Pressure Relief Valves -

-

-

Connection size and type: The inlet and discharge piping size must at least be as large as the inlet/ discharge opening on the valve itself. The connection type, for example, male/female flange is also important. All of these factors help to determine which valve to choose from the manufacturer’s catalog. Set pressure: The pressure differential at which a safety or pressure relief valve opens is its set pressure. This is the gauge pressure at which it is set on the test bench. Set pressure can be adjusted within limits by varying the spring tension that keeps the valve closed. Such calibration/ adjustments are done on the workshop bench before mounting the valve. The set pressure of the valve must be equal to or below the maximum allowable working pressure (MAWP) of the equipment it is supposed to protect. MAWP should be at least 10% more than the highest expected operating pressure under normal circumstances. In the case of multiple pressure relief/safety valves protecting the equipment, the second device set pressure may be 1.05 times or 1.10 times the MAWP according to the same code. Table 20.1 mentions the recommended pressure differentials as per ASME Section VIII, Div. I, Appendix M M-11. Temperature: Temperature is an important parameter as it affects the volume and viscosity of the flowing system. It is to be considered for selecting the suitable material of construction.

20.4 Blowdown and flare

739

- Backpressure: This is the pressure downstream of the valve when the valve is popping. It is desirable to have a fairly constant value of back pressure; else the popping pressure may be lower or higher than the set pressure. For installations with high levels of constant back pressure, a bellows-sealed valve or pilot-operated valve may be chosen. The required relief capacity may be decided by API 520/STD 521. The sizing of the orifice may be found using API RP 520 Part I and ASME Section I. Apart from the above, the designer also considers the physical dimensions of the equipment and the plant, as well as factors related to the environment in which the valve will operate. Table 20.1 Minimum recommended pressure differential (ASME Section VII, Div. I, Appendix M M-11). Set pressure (Ps)

Minimum differential pressure

Set pressure tolerance

Ps  483 kPa

35 kPa

13.8 kPa

483

FIGURE 21.4 (A) Diurnal inflow to ETP and (B) Cumulative inflow.

20

24

772

Chapter 21 Process packages

In Fig. 21.4B, along with the cumulative curve, a straight line corresponding to the average flow rate is drawn that has slope of 333.3 m3/hr. Tangents to the actual cumulative flow curve are drawn parallel to the mentioned straight line on both sides of the average flow line. The distance between the two lines, perpendicular to the x-axis, gives the minimum equalization volume required for the process. In this case, it is 1850 m3. For the segregated streams with 33% and 67% of total inflow, two equalization tanks of capacity (w620 and w1250 m3) are considered. The minimum submergence (S, m) to prevent strong air core vortex is estimated from a flowe diameter relationship: S ¼ d þ 0:574  Q=d 1:5 ; where S, D in inch, Q in US gpm. Pipe diameter:  • TDS stream: QTDS;WW ¼ 2 3Q ¼ ð2=3Þ  ð8000=24Þ ¼ 222:2 m3/hr ¼ 978.3 US gpm. 
p  d2 1:2 ¼ 222:2 3600 . Assumed velocity w1.2 m/s, which gives d ¼ 0.256 m 4 DN250 40 Sch. Pipe (di ¼ 254.46 mm ¼ 10.02 in.) is used and the revised velocity is obtained as U ¼ 1:2  ð0:256=0:25446Þ2 ¼ 1:215 m/s which is acceptable. This gives

 S ¼ 10:02 þ 0:574  978:3 ð10:02Þ1:5 ¼ 27:700 ¼ 705 mm. • Oily water stream: QOWW ¼ Q=3 ¼ ð1=3Þ  ð8000=24Þ ¼ 111:1 m3/hr ¼ 489.2 US gpm. 
p  d 2  1:2 ¼ 111:1 3600 . Assumed velocity w1.2 m/s, and d ¼ 0.181 m 4 DN200 40 sch. Pipe (di ¼ 202.74 mm ¼ 7.98 in.) is used, and revised velocity is U ¼ 1:2  ð0:181=0:20274Þ2 ¼ 0:956 m/s, which is acceptable. 0:574  489:2 The minimum submergence, S ¼ 7:98 þ ¼ 20:400 z525 mm. ð7:98Þ1:5 The two tanks need equalization volume of 620 and 1250 m3, respectively. Adding w10% safety allowance and making sure that there is an additional minimum submergence, the dimensions are selected such that the area used is minimum and the depth is as low as possible. This gives the dimensions of the two tanks as 11:7 m  11:7 m  5 m and 15:1 m  15:1 m  6 m. Design of DAF DAF clarifies WW by removing suspended matter such as oil or solids. The removal is achieved by introducing the gas phase directly into the liquid phase through a revolving impeller in a flotation tank or basin. The released air forms tiny bubbles which adhere to the suspended matter causing the same to float to the surface of the water from where it may be removed by a skimming device. To design a DAF tank, the air to solids ratio A=S needs to be known. It is somewhere around 0.005 to 0.06 for refinery WW treatment plants where ðA=SÞ is in mL air/mg solids, and the relationship is given by 1:3  sa  ðP  f  1Þ Si For the present problem, 0:02  A=S  0:03 is considered and ðA=SÞ ¼

sa ¼ air solubility, mL/L given as a function of temperature in Table 21.12. f ¼ fraction of air dissolved at pressure P, usually 0.5 P ¼ Pressure, atm Si ¼ influent suspended solids g/m3

21.2 Examples

773

Table 21.12 Air Solubility versus Temperature. Temperature,  C

0

10

20

30

sa , mL/L

29.2

22.8

18.7

15.7

Assuming the operating temperature to be 30 C (in Indian context), sa ¼ 15:7 mL/L and Si is considered as 400 mg/L. Thus, the pressure required will be P ¼ 2.78 e3.18 atm. for A/S ratio of 0.02 e 0.03. Design of activated sludge treatment Main components in an activated sludge treatment plant are pretreatment solid remover, primary clarifier, aeration tank and secondary clarifier. Sludge formed in this process goes to a thickener and sludge digester. The process schematic is shown in Fig. 21.5. Steps involved in designing an activated sludge treatment plant are as follows: i. ii. iii. iv. v. vi. vii. viii. ix.

Design solid removal equipment Determine BOD5 of effluent Determine treatment efficiency Determine reactor volume Determine HRT Determine return sludge flow rate Determine surplus sludge production Determine organic loading Design clarifiers

Assumptions i. ii. iii. iv. v.

Temperature w30 C. Ratio of volatile suspended solid to total suspended solid, VSS=TSS ¼ 0.8. Return sludge concentration ¼ 10,000 mg/L. MLSS ¼ 3500 mg/L in aeration tank. Design mean cell residence time, qc ¼ 10 days.

Solid Mixed stream removal

Primary clarifier

Aeration basin

Clarifier/ Settler

Sand filter

Air Sludge

FIGURE 21.5 Schematic of activated sludge treatment.

Treated effluent

774

vi. viii. x. xi.

Chapter 21 Process packages

Effluent contains 20 mg/L (10% TSS) of biological solids, 60% of VSS is biodegradable. WW contains adequate nitrogen, phosphorus and trace nutrients for biological growth. 30% raw BOD5 is removed in primary sedimentation. Incoming total Kjeldahl nitrogen (TKN) ¼ 53 mg/L

Known parameters i. Flow rate Q ¼ 8000 m3/d ¼ w0.1 m3/s ii. BOD5 ¼ 600 mg/L (some BOD is also removed by DAF), BOD5 @ basin inlet ¼ 0:7  600 ¼ 420 mg/L iii. Final effluent BOD5 target ¼ w17 mg/L iv. Yield coefficient Y ¼ 0.6 v. Decay constant kd ¼ 0:07 per day mg=L vi. Specific substrate utilization rate q ¼ 0:038 day Determination of effluent BOD5   1 1  Substrate concentration: S ¼ þ kd ¼ 7.45 mg/L. qy qc Total BOD5 Effluent TSS ¼ 20 mg/L, VSS/SS ¼ 0.8, degradable fraction of VSS ¼ 0.6. BOD5 of VSS ¼ 0:6  0:8  20 ¼ 9:6 mg/L. Total BOD5 ¼ 9.6 þ 7.45 ¼ 17.05 mg/L. ðS0  SÞ ð600  17:05Þ  100 ¼ 97:16%  100 ¼ Efficiency, E ¼ S0 600 Effluent TSS ¼ 20 mg/L and given that X¼

VSS SS

¼ 0.8,

VSS  MLSS ¼ 0:8  3500 ¼ 2800 SS

qc Q  Y  ðS0  SÞ  1 þ kd  qc , here S0 is 70% of BOD, because 30% is removed in sedimentation, i.e., X S0 ¼ 0:7  600 ¼ 420, and

V ¼

10 8000  0:6  ð420  17:05Þ  ¼ 4063 ¼ w4100 m3 2800 1 þ 0:07  10 Dimension: 7 m  10:5 m  60 m (¼w4400 m3 > 4100 m3, with sufficient margin) HRT ¼ ð7  10:5  60Þ=8000 ¼ 0:551 day V ¼

Return sludge flow rate MLSS 3500 ¼ ¼ 0:5384 Recycle ratio RR ¼ RSC  MLSS 10; 000  3500 Recycle flow rate ¼ QR ¼ RR  Q ¼ 0:384  8000 ¼ 4307 m3/day. Surplus sludge to be disposed off r ¼ 1000 g/L VSS ¼

V  X 103 4307  2800 103  ¼  ¼ 1:21 mt=day qc 10 r 103 SS ¼ 1.21=0.8 ¼ 1.51 mt=day

Further reading

775

Further reading Coulson, J. M., & Warner, F. E. (1949). A problem in chemical engineering design: The Manufacture of Mononitrotoluene. Great Britain): Institution of Chemical Engineers. Peters, M. S., & Timmerhaus, K. D. (1980). Plant design and economics for chemical engineers. McGraw-Hill.

Appendix A: Graphical symbols for piping systems and plant Based on BS 1553: PART 1: 1977 Scope This part of BS 1553 specifies graphical symbols for use in flow and piping diagrams for process plant.

Symbols (or elements of Symbols) for Use in Conjunction with Other Symbols Mechanical linkage

Weight device

Electrical device

Access point

Equipment branch: general symbol. Note: The upper representation does not necessarily imply a flange, merely the termination point. Where a breakable connection is required, the branch/pipe would be as shown in the lower symbol.

Vibratory or loading device (any type)

Equipment penetration (fixed)

Spray device

Equipment penetration (removable)

Rotary movement

Boundary line

Stirring device

Point of change

Fan

Discharge to atmosphere

Reprinted from Ray Sinnott and Gavin Towler, Appendix A - Graphical symbols for piping systems and plant.

777

778

Appendix A: Graphical symbols for piping systems and plant

Basic and Developed Symbols for Plant and Equipment Heat Transfer Equipment Heat exchanger (basic symbols)

Alternative: Shell and tube: fixed tube sheet

Shell and tube: U tube or floating head

Shell and tube: kettle reboiler

Air-blown cooler

Plate type

Double pipe type

Heating / cooling coil (basic symbol)

Fired heater / boiler (basic symbol)

Appendix A: Graphical symbols for piping systems and plant

Upshot heater

Detail A Where complex burners are employed the ‘‘burner block’’ may be detailed elsewhere on the drawing.

Detail A

Vessels and Tanks Drum or simple pressure vessel (basic symbol)

Knock-out drum (with demister pad)

Tray column (basic symbol)

Tray column: Trays should be numbered from the bottom; at least the first and the last should be shown. Intermediate trays should be included and numbered where they are significant.

30

14

779

780

Appendix A: Graphical symbols for piping systems and plant

Fluid contacting vessel (basic symbol)

Fluid contacting vessel: Support grids and distribution details may be shown

Reaction or absorption vessel (basic symbol)

Reaction or absorption vessel: Where it is necessary to show more than one layer of material, alternative hatching should be used

Autoclave (basic symbol)

Autoclave

Appendix A: Graphical symbols for piping systems and plant

Open tank (basic symbol)

Open tank

Clarifier or settling tank

Sealed tank

Covered tank

Tank with fixed roof (with draw-off sump)

Tank with floating roof (with roof drain)

Storage sphere

Gas holder (basic symbol for all types)

781

782

Appendix A: Graphical symbols for piping systems and plant

Pumps and Compressors Rotary pump, fan, or simple compressor (basic symbol)

Centrifugal pump or centrifugal fan

Centrifugal pump (submerged suction)

Positive displacement rotary pump or rotary compressor

Positive displacement pump (reciprocating)

Axial flow fan

Compressor: centrifugal/axial flow (basic symbol)

Compressor: centrifugal/axial flow

Compressor: reciprocating (basic symbol)

Ejector/injector (basic symbol)

Appendix A: Graphical symbols for piping systems and plant

Solids Handling Size reduction

Breaker gyratory

Roll crusher

Pulverizer: ball mill

Mixing (basic symbol)

Kneader

Ribbon blender

Double cone blender

Filter (basic symbol, simple batch)

Filter press (basic symbol)

Rotary filter, film, drier or flaker

783

784

Appendix A: Graphical symbols for piping systems and plant

Cyclone and hydroclone (basic symbol)

Cyclone and hydroclone

Centrifuge (basic symbol)

Centrifuge: horizontal peeler type

Centrifuge: disc bowl type

Drying Drying oven

Belt drier (basic symbol)

Rotary drier (basic symbol)

Rotary kiln

Appendix A: Graphical symbols for piping systems and plant

Spray drier

Belt conveyor

Screw conveyor

Elevator (basic symbol)

Electric motor (basic symbol)

Turbine (basic symbol)

785

Appendix B: Corrosion chart An R indicates that the material is resistant to the named chemical up to the temperature shown, subject to the limitations given in the notes. The notes are given at the end of the table. A blank indicates that the material is unsuitable. ND indicates that no data was available for the particular combination of material and chemical. This chart is reproduced with the permission of IPC Industrial Press Ltd.

Note

This appendix should be used as a guide only. Before a material is used, its suitability should be cross-checked with the manufacturer.

Appendix B - Corrosion Chart In Chemical Engineering Design 6E, SI Edition, © 2020, with permission from Elsevier.

787

788

Appendix B: Corrosion chart

Appendix B: Corrosion chart

789

790

Appendix B: Corrosion chart

Appendix B: Corrosion chart

791

792

Appendix B: Corrosion chart

Appendix B: Corrosion chart

793

794

Appendix B: Corrosion chart

Appendix B: Corrosion chart

795

796

Appendix B: Corrosion chart

Appendix B: Corrosion chart

797

798

Appendix B: Corrosion chart

Appendix B: Corrosion chart

799

800

Appendix B: Corrosion chart

Appendix B: Corrosion chart

801

802

Appendix B: Corrosion chart

Appendix B: Corrosion chart

803

804

Appendix B: Corrosion chart

Appendix B: Corrosion chart

805

Appendix C: Physical property data bank Inorganic compounds are listed in alphabetical order of the principal element in the empirical formula. Organic compounds with the same number of carbon atoms are grouped together and arranged in order of the number of hydrogen atoms, with other atoms in alphabetical order. A searchable spreadsheet containing the physical property data and models is available in the online material at http://books.elsevier.com/companions. NO ¼ Number in list MOLWT ¼ Molecular weight TFP ¼ Normal freezing point, deg C TBP ¼ Normal boiling point, deg C TC ¼ Critical temperature, deg K PC ¼ Critical pressure, bar VC ¼ Critical volume, cubic meter/mol LDEN ¼ Liquid density, kg/cubic meter TDEN ¼ Reference temperature for liquid density, deg C HVAP ¼ Heat of vaporization at normal boiling point, J/mol VISA, VISB ¼ Constants in the liquid viscosity equation: LOG[viscosity] ¼ [VISA] * [(1/T)  (1/VISB)], viscosity mNs/sq.m, T deg K DELHF ¼ Standard enthalpy of formation of vapor at 298 K, kJ/mol. DELGF ¼ Standard Gibbs energy of formation of vapor at 298 K, kJ/mol. CPVAPA, CPVAPB, CPVAPC, CPVAPD ¼ Constants in the ideal gas heat capacity equation: Cp ¼ CPVAPA + (CPVAPB) *T + (CPVAPC) *T **2 + (CPVAPD) *T **3, Cp ¼ J/mol K, T deg K ANTA, ANTB, ANTC ¼ Constants in the Antione equation: Ln (vapor pressure) ¼ ANTA e ANTB/(T + ANTC), vapour pressure mmHg, T deg K To convert mmHg to N/sq.m, multiply by 133.32. To convert degrees Celsius to Kelvin, add 273.15. TMN ¼ Minimum temperature for Antoine constant, deg C. TMX ¼ Maximum temperature for Antoine constant, deg C Most of the values in this data bank were taken, with the permission of the publishers, from The Properties of Gases and Liquids, third ed., by R. C. Reid, T. K. Sherwood, and J. M. Prausnitz, McGraw-Hill.

807

808

Appendix C: Physical property data bank

Appendix C: Physical property data bank

809

810

Appendix C: Physical property data bank

Appendix C: Physical property data bank

811

812

Appendix C: Physical property data bank

Appendix C: Physical property data bank

813

814

Appendix C: Physical property data bank

Appendix C: Physical property data bank

815

816

Appendix C: Physical property data bank

Appendix C: Physical property data bank

817

818

Appendix C: Physical property data bank

Appendix C: Physical property data bank

819

820

Appendix C: Physical property data bank

Appendix C: Physical property data bank

821

822

Appendix C: Physical property data bank

Appendix C: Physical property data bank

823

824

Appendix C: Physical property data bank

Appendix C: Physical property data bank

825

826

Appendix C: Physical property data bank

Appendix C: Physical property data bank

827

Appendix D: Conversion factors Item Length

Time

Area

Volume

Volumetric flow rate

Unit 1 1 inch

¼

0.0254 m

1 ft

¼

0.3048 m

1 min

¼

60 s

1 hr

¼

3600 s

1 day

¼

86400 s

1 year (annum) of plant operation

¼

8000 or 8200 hr nominally

1 ft2

¼

0.092903 m2

1 acre

¼

4046.86 m2

1 hectare

¼

10000 m2

1 ft3

¼

28.3168 litre

1 litre

¼

0.001 m3

1 US gallon

¼

3.78541 litre

1 Imperial gallon

¼

4.54609 litre

1 barrel

¼

42 US gallon

3

1 ft /min

¼

28.3168 litre/min

1 ft3/s

¼

28.3168 litre/s

1 m /hr

¼

1000 litre/hr

1 mt/d

¼

0.01175979259 kg/s

3

Mass

Mass flow rate

Unit 2

1 (US) gpm

¼

0.06309019667 litre/s

1 (Imperial) gpm

¼

0.0757682 litre/s

1 lb

¼

0.4536 kg

1 (long - UK) ton

¼

2240 lb

1 (long - UK) ton

¼

1016.047 kg

1 (short - US) ton

¼

2000 lb

1 (short - US) ton

¼

907.1847 kg

1 metric ton (mt) or ton

¼

1000 kg

1 lb/hr

¼

0.000125998 kg/s

1 mt/hr

¼

0.277778 kg/s Continued

829

830

Appendix D: Conversion factors

Force

Temperature difference

Energy (Work, Heat)

1 kgf

¼

9.80665 N

1 lbf

¼

4.44822 N

1 dyne

¼

105 N

1 oF

¼

5

1 oR

¼

5

1 oC

¼

1 oK

1 cal

¼

4.1868 J

1 erg

¼

107 J

1 BTU

¼

1.05506 kJ

1 BTU

¼

252 cal

1 hp.hr

¼

2.6845 MJ

1 kwh

/9 oC /9 oC

¼

3.6 MJ

2

¼

4.8824 kg/m2

1 lb/in2

¼

703.07 kg/m2

1 lb/ft3

¼

16.019 kg/m3

1 lb/(US gallon)

¼

119.83 kg/m3

1 lb/(Imperial gallon)

¼

99.776 kg/m3

1 psi

¼

6.8948 kPa

1 standard atm

¼

101.325 kPa

1 atm

¼

98.0667 kPa

1 bar

¼

102 kPa

1 ft water

¼

2.9891 kPa

1 in water

¼

249.09 Pa

1 Pa

¼

1 N/m2

1 in Hg

¼

3.3864 kPa

1 mm Hg (1 torr)

¼

133.32 Pa

1 hp (British)

¼

745.7 W

1 hp (Metric)

¼

735.5 W

1 BTU/hr

¼

0.29307 W

1 ton of refrigeration

¼

3516.9 W

Viscosity, dynamic

1 Poise

¼

0.1 Pa.s

1 lb/(ft.hr)

¼

0.41338 mPa.s

Viscosity, kinematic

1 cSt (centiStokes)

¼

102 m2/s

2

1 ft /hr

¼

0.25806 cm2/s

1 ft2/hr

¼

25.806 x 106 m2/s

Mass per unit area

Density

Pressure

Power

1 lb/ft

Appendix D: Conversion factors

Surface tension

1 dyne/cm

Mass flux

1 lb/(ft2.hr)

¼

103 N/m

¼

0.0013562 kg/(m2.s)

Heat transfer coefficient

1 BTU/(hr.ft .F)

¼

5.6783 W/(m2.K)

Latent heat

1 BTU/lb

¼

2.326 kJ/kg

1BTU/(lb. F)

¼

4.1868 kJ/(kg.K)

o

¼

1.163 W/(m.K)

Specific heat Thermal conductivity

2

o

1 kcal/(hr.m. C)

831

Appendix E: Typical fouling factors in m2K/W compiled from various sources Table E.1A: Fouling factors e cooling water services Water temperature

Cooling water 1 m/s

1 m/s

2

Type of water

Fouling factors in [m K/W]

Sea

0.00009

0.00009

0.00018

0.00018

Brackish

0.00035

0.00018

0.00053

0.00035

Cooling tower

0.00018

0.00018

0.00035

0.00035

Cooling tower untreated

0.00053

0.00053

0.00088

0.00070

Tap water

0.00018

0.00018

0.00035

0.00035

River minimum

0.00018

0.00018

0.00035

0.00035

River average

0.00053

0.00035

0.00070

0.00035

Engine jacket

0.00018

0.00018

0.00018

0.00018

DM water

0.00009

0.00009

0.00009

0.00009

Treated boiler feedwater

0.00018

0.00009

0.00018

0.00018

Boiler blowdown

0.00035

0.00035

0.00035

0.00035

833

834

Appendix E: Typical fouling factors in m2K/W compiled from various sources

Table E.1B: Fouling factors: miscellaneous gases, vapour and liquids Fouling Factors in m2K/W Fluid Oil

Gases

Liquids

Gas and vapour

Fouling factor [m2K/W]

Gas oil

0.00009

Transformer oil/heat transfer oil/hydraulic oil

0.00018

Hydrogen

0.00176

Engine exhaust

0.00176

Steam

0.00009

Steam with oil traces

0.00018

Cooling fluid vapours with oil traces

0.00035

Organic solvent vapours

0.00018

Compressed air

0.00035

Natural gas

0.00018

Stable top products

0.00018

Cooling fluid

0.00018

Organic heat transfer fluids

0.00018

Salts

0.00009

LPG, LNG

0.00018

MEA and DEA (amines) solutions

0.00035

DEG and TEG (glycols) solutions

0.00035

Stable side products

0.00018

Stable bottom products

0.00018

Caustics

0.00035

Vegetable oils

0.00053

Hydrogen

0.00176

Engine exhaust

0.00176

Steam

0.00009

Steam with oil traces

0.00018

Cooling fluid vapours with oil traces

0.00035

Appendix E: Typical fouling factors in m2K/W compiled from various sources

Organic solvent vapours

0.00018

Compressed air

0.00035

Natural gas

0.00018

Stable top products

0.00018

835

Table E.1C: Fouling factors: refinery streams 0e95 C Temperature Crude e wet (contains brine)

Crude e after desalting

Velocity

Velocity

Crude distillation streams

Cracking and coking unit streams

95e160 C

160e260 C

>95 C

2

Fouling factor [m K/W]

< 0.6 m/s

0.00053

0.00088

0.00105

0.00123

0.6 m/s to 1.2 m/s

0.00035

0.00070

0.00088

0.00106

> 1.2 m/s

0.00035

0.00070

0.00070

0.00088

< 0.6 m/s

0.00053

0.00053

0.00070

0.00088

0.6 m/s to 1.2 m/s

0.00035

0.00035

0.00053

0.00070

> 1.2 m/s

0.00035

0.00035

0.00035

0.00053

Gasoline

0.00018

Naphtha

0.00018

Kerosene/ATF

0.00018

Light gas oil

0.00035

Heavy gas oil

0.00053

Reduced crude oil

0.00176

Overhead vapour

0.00035

Light cycle oil

0.00035

Heavy cycle oil

0.00053

Coker gas oil

0.00053

Heavy coker gas oil

0.00070

FCC bottom slurry (>1.5 m/s)

0.00053

Bottom liquid products

0.00035 Continued

836

Appendix E: Typical fouling factors in m2K/W compiled from various sources

0e95 C Temperature Naphtha catalytic reforming, hydrocracking, hydrodesulfurization (HDS) streams

Light ends processing streams

Lube refinery streams

95e160 C

160e260 C

Fouling factor [m2K/W]

Reformer feed

0.00035

Reformate

0.00018

Hydrocracker charge and effluent

0.00035

Recycle gas

0.00018

HDS feed and effluent

0.00035

Overhead vapour

0.00018

Liquid stream above 50 API

0.00018

Liquid stream between 30 and 50 API

0.00035

Overhead vapour

0.00018

Liquid products

0.00018

Absorption oils

0.00035

Alkylation trace acid streams

0.00053

Reboiler streams

0.00053

Lube oil processing streams

0.00053

Feed stock

0.00035

Solvent feed mix

0.00035

Solvent

0.00018

Extract

0.00035

Raffinate

0.00018

Asphalt

0.00088

Wax slurries

0.00053

Finished lube oil

0.00018

>95 C

Appendix F: Heat exchanger tube sizes and other details Table F.1A: Tubes of different materials (TEMA) Materials Carbon steel

Tube side fouling factor Less than 0.00035 m2K/W

Low alloy steel (