Desalination [2nd edition.] 9781119407744

595 15 12MB

English Pages [755] Year 2019

Report DMCA / Copyright

DOWNLOAD FILE

Polecaj historie

Desalination [2nd edition.]
 9781119407744

Citation preview

Desalination 2nd Edition

Scrivener Publishing 100 Cummings Center, Suite 541J Beverly, MA 01915-6106 Publishers at Scrivener Martin Scrivener ([email protected]) Phillip Carmical ([email protected])

Desalination 2nd Edition

Water from Water

Edited by

Jane Kucera

This edition first published 2019 by John Wiley & Sons, Inc., 111 River Street, Hoboken, NJ 07030, USA and Scrivener Publishing LLC, 100 Cummings Center, Suite 541J, Beverly, MA 01915, USA © 2019 Scrivener Publishing LLC For more information about Scrivener publications please visit www.scrivenerpublishing.com. All rights reserved. No part of this publication may be reproduced, stored in a retrieval system, or transmitted, in any form or by any means, electronic, mechanical, photocopying, recording, or otherwise, except as permitted by law. Advice on how to obtain permission to reuse material from this title is available at http://www.wiley.com/go/permissions. Wiley Global Headquarters 111 River Street, Hoboken, NJ 07030, USA For details of our global editorial offices, customer services, and more information about Wiley products visit us at www.wiley.com. Limit of Liability/Disclaimer of Warranty While the publisher and authors have used their best efforts in preparing this work, they make no representations or warranties with respect to the accuracy or completeness of the contents of this work and specifically disclaim all warranties, including without limitation any implied warranties of merchantability or fitness for a particular purpose. No warranty may be created or extended by sales representatives, written sales materials, or promotional statements for this work. The fact that an organization, website, or product is referred to in this work as a citation and/or potential source of further information does not mean that the publisher and authors endorse the information or services the organization, website, or product may provide or recommendations it may make. This work is sold with the understanding that the publisher is not engaged in rendering professional services. The advice and strategies contained herein may not be suitable for your situation. You should consult with a specialist where appropriate. Neither the publisher nor authors shall be liable for any loss of profit or any other commercial damages, including but not limited to special, incidental, consequential, or other damages. Further, readers should be aware that websites listed in this work may have changed or disappeared between when this work was written and when it is read. Library of Congress Cataloging-in-Publication Data ISBN 978-1-119-40774-4 Background Image (desalination plant): Станислав Саблин | Dreamstime.com Cover design by Kris Hackerott Set in size of 11pt and Minion Pro by Exeter Premedia Services Private Ltd., Chennai, India Printed in the USA 10 9 8 7 6 5 4 3 2 1

DEDICATION: In fond memory of , Julius “Bud” Glater my thesis advisor and mentor, and a pioneer in the development of membrane-based, desalination technologies.

Contents Preface 1 Introduction to Desalination Jane Kucera 1.1 Introduction 1.2 How Much Water is There? 1.2.1 Global Water Availability 1.2.2 Water Demand 1.2.3 Additional Water Stress Due to Climate Change 1.3 Finding More Fresh Water 1.3.1 Relocating Water 1.3.2 Conservation and Reuse 1.3.3 Develop New Sources of Fresh Water 1.4 Desalination: Water from Water 1.4.1 Drivers for Desalination 1.4.2 Feed Water Sources for Desalination 1.4.3 Current Users of Desalinated Water 1.4.4 Overview of Desalination Technologies 1.4.5 History of Desalination Technologies 1.4.5.1 History of Thermal Desalination 1.4.5.2 History of Reverse Osmosis Desalination 1.4.5.3 Developments in Desalination Since 1980 1.4.6 The Future of Desalination 1.5 Desalination: Water from Water Outline Abbreviations References 2

Thermal Desalination Processes 2.1 Introduction 2.2 Mass- and Energy Balances 2.2.1 Single-Stage Evaporation 2.2.2 Multiple-Effect Evaporation 2.2.3 Multi-Stage-Flash Evaporation

xxi 1 1 2 2 4 6 7 7 9 11 15 15 16 20 21 24 24 26 27 34 41 43 44 51 51 52 52 61 80 vii

viii

Contents 2.2.4

Multiple-Effect Distillation with Thermal Vapour Compression (MED-TVC) 2.2.5 Single-Stage Evaporation with Mechanically Driven Vapour Compression 2.3 Performance of Thermal Desalination Processes 2.3.1 Definition of Gained Output Ratio 2.3.2 Single Purpose vs. Dual Purpose Plants 2.3.3 Specific Primary Energy Consumption 2.4 Recent Developments in Thermal Desalination Processes 2.4.1 Hybrid Plants 2.4.1.1 Multi-Stage Flash with Reverse Osmosis (MSF-RO) 2.4.1.2 Multi-Effect Distillation with Reverse Osmosis (MED-RO) 2.4.2 Expanding the Scope of Hybrid Thermal Desalination Future Prospects 2.5 2.5.1 General Remarks 2.5.2 Optimization of Existing Process Design 2.5.2.1 Material of Construction 2.5.2.2 Increasing Water Velocity 2.5.2.3 Heat Transfer Enhancement by Using Corrugated Oval Tubes 2.5.2.4 Increasing the Top Operation Temperature to 85 °C 2.5.2.5 Increasing Number of Stages 2.5.2.6 Modifications in MED-TVC References 3 Basic Terms and Definitions 3.1 Reverse Osmosis System Flow Rating 3.2 Recovery 3.3 Rejection 3.4 Flux 3.5 Concentration Polarization 3.6 Beta 3.7 Fouling 3.8 Scaling 3.9 Silt Density Index 3.10 Modified Fouling Index 3.11 Langelier Saturation Index References

92 104 111 111 114 126 131 131 131 132 132 134 134 134 134 135 136 136 136 136 137 139 139 140 142 145 147 148 149 152 154 157 160 161

Contents ix 4 Nanofiltration – Theory and Application Christopher Bellona 4.1 Introduction 4.2 Defining Nanofiltration 4.3 History of Nanofiltration 4.4 Theory 4.4.1 Mechanisms of Solute Removal 4.4.1.1 Ion Rejection 4.4.1.2 Organic Solute Rejection 4.4.2 Modeling NF Separations 4.4.2.1 Donnan Steric Pore Model 4.4.2.2 Irreversible Thermodynamic or Phenomenological Model 4.4.2.3 Other Modeling Approaches 4.4.3 Membrane Fouling 4.5 Application 4.5.1 Water and Wastewater Treatment Industry 4.5.1.1 Water Treatment 4.5.1.2 Wastewater Treatment and Reuse 4.5.1.3 Desalination 4.5.2 Food Industry 4.5.2.1 Dairy Industry 4.5.2.2 Sugar and Beverage Industry 4.5.3 Chemical Processing Industry 4.5.3.1 Pharmaceutical Industry 4.5.3.2 Textile Industry 4.6 Conclusions References

163

5 Forward Osmosis Jeffrey McCutcheon, Lingling Xia and Nhu-Ngoc Bui 5.1 The Limitations of Conventional Desalination 5.1.1 Osmotic Pressure 5.2 Forward Osmosis 5.2.1 History of FO 5.2.2 Benefits of Forward Osmosis 5.3 The Draw Solution 5.3.1 Inorganic Solutes 5.3.2 Nanomaterials 5.3.3 Organic Solutes 5.4 The Membrane

209

163 164 168 170 171 171 173 177 177 178 179 180 182 182 182 183 187 189 189 191 191 192 192 193 194

210 210 212 212 214 215 216 217 217 218

x

Contents 5.4.1 Mass Transfer Limitations in Forward Osmosis 5.4.2 Tailored Membranes for FO 5.4.2.1 Flat Sheet 5.4.2.2 Hollow Fiber 5.5 Process Design and Desalination Applications 5.6 Future Directions Acknowledgements References

219 221 222 224 226 232 234 234

6 Electrodialysis Desalination Jae-Hwan Choi, Hong-Joo Lee and Seung-Hyeon Moon 6.1 Principles of ED 6.2 Preparation and Characterization of Ion Exchange Membranes 6.2.1 Preparation of Ion Exchange Membranes 6.2.2 Characterization of Ion Exchange Membranes 6.2.3 Concentration Polarization and the Limiting Current Density 6.3 ED Equipment Design and Desalination Process 6.3.1 ED Stack Design 6.3.2 ED Process Design 6.3.3 ED Operation and Maintenance 6.3.4 Design Parameters in Desalting ED 6.3.5 Economics of the ED Process 6.4 Control of Fouling in an ED Desalination Process 6.4.1 Fouling Mechanism 6.4.2 Fouling Potential 6.4.3 Fouling Mitigation 6.5 Prospects for ED Desalination 6.5.1 Integration with ED for the Desalination 6.5.2 Process Intensification of the ED Desalination System 6.5.3 ED Powered by Photovoltaic Solar Energy 6.5.4 Perspectives of ED Desalination 6.6 Concluding Remarks References

245

7 Continuous Electrodeionization Jonathan H. Wood and Joseph D. Gifford 7.1 Introduction 7.2 Development History

287

246 249 249 251 253 261 261 262 264 265 267 270 270 271 273 275 275 276 278 280 281 282

287 289

Contents xi 7.3 Technology Overview 7.3.1 Mechanisms of Ion Removal 7.4 CEDI Module Construction 7.4.1 Device Configurations 7.4.2 Resin Configurations 7.4.2.1 Mixed Bed Resin Filler (CEDI-MB) Intermembrane Spacing 7.4.2.2 Mixed Bed Resin Filler (CEDI-MB) Resin Packing 7.4.2.3 Layered Bed Resin Filler (CEDI-LB) 7.4.2.4 Separate Bed Resin Filler (CEDI-SB) 7.4.3 Flow Spacers 7.5 Electroactive Media Used in CEDI Devices 7.5.1 Ion Exchange Resin Selection 7.5.2 Ion Exchange Membrane Selection 7.6 DC Current and Voltage 7.6.1 Faraday’s Law 7.6.2 Current Efficiency and E-Factor 7.6.3 Ohm’s Law and Module Resistance 7.6.4 Electrode Reactions and Material Selection 7.7 System Design Considerations 7.7.1 Required Process Control & Instrumentation 7.7.2 Optional Process Control & Instrumentation 7.8 Process Design Considerations 7.8.1 Feed Water Requirements 7.8.2 Hardness 7.8.3 Carbon Dioxide 7.8.4 Oxidants 7.8.5 Temperature 7.8.6 Water Recovery 7.8.7 Recycling of CEDI Reject Stream 7.8.8 Total Organic Carbon 7.8.9 Electrode Gases 7.9 Operation and Maintenance 7.9.1 Estimation of Operating Current and Voltage 7.9.2 Power Supply Operation 7.9.3 Power Consumption 7.9.4 Flows and Pressures 7.9.5 Record Keeping 7.9.6 Cleaning and Sanitization 7.9.7 Preventive Maintenance

289 291 291 291 293 293 294 294 296 298 299 299 299 300 300 301 302 303 304 304 305 306 307 308 309 310 311 312 313 313 314 316 316 316 317 317 319 319 322

xii

Contents 7.10

Applications 7.10.1 Pharmaceutical and Biotechnology 7.10.2 Steam Generation 7.10.3 Microelectronics/Semiconductor 7.10.4 System Sizing 7.11 Future Trends Nomenclature References

322 322 323 323 324 324 325 326

8 Membrane Distillation: Now and Future Xing Yang, Anthony G. Fane and Rong Wang 8.1 Introduction 8.2 MD Concepts and Historic Development 8.2.1 MD Concepts and Configurations 8.2.2 Historic Development 8.3 MD Transport Mechanisms 8.3.1 Mass Transfer in MD 8.3.1.1 Mass Transfer Through the Feed Boundary Layer (CP Effect) 8.3.1.2 Mass Transfer Through Membrane Pores 8.3.2 Heat Transfer in MD 8.3.2.1 Heat Transfer on the Feed Side (TP Effect) 8.3.2.2 Heat Transfer Across the Membrane— Conduction and Evaporation 8.4 Strategic Development for an Enhanced MD System 8.4.1 MD Membranes 8.4.2 MD Module Design 8.4.3 MD Process Parameters 8.5 Energy and Cost Evaluation in MD 8.5.1 Thermal Efficiency and Cost Evaluation 8.5.2 Current Status of MD Cost and Energy Resources 8.6 Innovations on MD Application Development 8.7 Concluding Remarks and Future Prospects References

329

9 Humidification Dehumidification Desalination John H. Lienhard V 9.1 Introduction 9.1.1 Classification of HDH cycles 9.1.2 System-Level Performance Parameters 9.1.3 Improving the Energy Efficiency of HDH Systems 9.1.4 Components of the HDH System

387

329 331 331 334 336 337 337 338 341 342 342 343 343 351 356 358 359 362 364 367 370

387 390 391 394 395

Contents xiii 9.2

Thermal Design 9.2.1 Effectiveness Model (On-Design Model) 9.2.1.1 Water Heated HDH Cycle 9.2.1.2 Single and Multi-Stage Air Heated Cycle 9.2.1.3 Varied Pressure Cycles and Other Carrier Gases 9.2.1.4 Summary of On-Design Findings 9.2.2 Single-Stage Fixed-Area HDH (Off-Design model) 9.2.2.1 Optimal Performance of a Single-Stage System 9.2.2.2 Relationship of HCRd = 1 to Entropy Generation Minimization 9.2.2.3 Variation of GOR with Top Temperature 9.2.2.4 Summary of Off-Design Findings 9.3 Systems with Mass Extraction and Injection 9.3.1 System Balancing Algorithms (On-Design Model) 9.3.2 Balancing Fixed-Area Systems by Extraction/Injection (Off Design Analysis) 9.3.3 Experimental Realization of HDH with and without Extraction/Injection 9.3.4 Summary of HDH Characteristics Related to Extraction/Injection 9.4 Bubble Column Dehumidification 9.4.1 Modeling and Experimental Validation 9.4.2 Multistage Bubble Column Dehumidifiers 9.4.3 Coil-Free Bubble Columns 9.5 Effect of High Salinity Feed on HDH Performance Acknowledgments Nomenclature References 10 Freezing-Melting Desalination Processes Mohammad Shafiur Rahman and Mohamed Al-Khusaibi 10.1 Introduction 10.2 Background or History of Freezing-Melting Process 10.3 Principles of Freezing-Melting Process 10.4 Major Types of Freezing-Melting Process 10.5 Direct-Contact Freezing 10.5.1 Ice Nucleation 10.5.1.1 Ice-Crystallization Unit 10.5.1.2 Hydrate Formation

396 398 399 406 407 408 408 409 410 413 416 416 419 421 422 425 426 428 428 431 433 437 437 439 433 447 448 450 451 451 451 452 456

xiv

Contents 10.5.2 Ice Separation Unit 10.5.3 Wash Columns 10.5.3.1 Melting Unit 10.6 Gas Hydrate Process 10.7 Direct-Contact Eutectic Freezing 10.8 Indirect-Contact FM Process 10.8.1 Internally Cooled 10.8.1.1 Progressive Static Layer Growth System as Block of Ice 10.8.1.2 Progressive Dynamic Layer Growth (Falling Film Type) 10.8.1.3 Progressive Dynamic Layer Growth (Circular Tube Type) 10.8.1.4 Melting of Progressive Layer or Block Crystals 10.8.1.5 Progressive Layer Crystallization on Rotating Drum 10.8.1.6 Progressive Suspension Growth 10.8.2 Externally Cooled 10.9 Pressure and Vacuum Processes 10.9.1 Vacuum System 10.9.2 Vapor-Compression System 10.9.3 Vapor-Absorption 10.9.4 Multiple-Phase Transformation 10.9.5 Pressure-Shift Nucleation and FM Process 10.10 Applications 10.11 Future Challenges Acknowledgment Abbreviations References

11 Ion Exchange in Desalination Bill Bornak 11.1 Introduction 11.2 Early Ion Exchange Desalination Processes 11.3 Life After RO 11.4 Ion Exchange Softening as Pre-Treatment 11.5 Softening by Ion Exchange 11.6 Boron-Selective Ion Exchange Resins as Post-Treatment 11.7 New Vessel Designs

457 457 459 459 459 460 460 460 461 462 462 463 463 464 464 464 465 465 465 466 466 469 470 471 471 479 480 480 482 483 485 486 491

Contents 11.8 New Resin Bead Design 11.9 Conclusion References 12 Electrosorption of Heavy Metals with Capacitive Deionization: Water Reuse, Desalination and Resources Recovery Pei Xu, Brian Elson and Jörg E Drewes 12.1 Introduction 12.1.1 Removal of Heavy Metals from Aqueous Solutions 12.1.2 Capacitive Deionization 12.2 Experimental Methods 12.2.1 CDI Treatment System 12.2.2 Feed Water Quality and Sample Analysis 12.3 Results and Discussions 12.3.1 CDI Voltage and Current Profiles 12.3.2 Removal of Heavy Metals from Electrolytes 12.3.3 Removal of Cyanide 12.4 Conclusions References 13 Solar Desalination Eydhah Almatrafi, D. Yogi Goswami, Mohammad Abutayeh, Chennan Li and Elias K. Stefanakos 13.1 Introduction 13.2 Solar Desalination 13.2.1 Conventional Desalination 13.2.2 Renewable Energy Driven Desalination 13.2.3 Solar Energy-Driven Desalination 13.3 Direct Solar Desalination 13.3.1 Solar Still 13.3.2 Solar-Driven Humidification– Dehumidification (HDH) 13.4 Indirect Solar Desalination 13.4.1 Phase Change Processes 13.4.1.1 Solar-Assisted Multi-Stage Flash 13.4.1.2 Solar-Assisted MultipleEffect Distillation 13.4.1.3 Solar-Assisted Heat Pumps (HP)

xv 493 494 495

497 498 498 500 502 502 504 506 506 507 514 518 516 525

526 528 528 528 529 530 530 532 533 533 534 536 541

xvi

Contents 13.4.2 Membrane Processes 13.4.2.1 Solar-Driven Reverse Osmosis 13.4.2.2 Solar-Driven Electro-Dialysis 13.4.2.3 Solar Thermal Driven Membrane Distillation (MD) 13.5 Non-Conventional Solar Desalination 13.5.1 Solar-Assisted Passive Vacuum 13.5.2 Power−Water Cogeneration 13.6 Solar Integration and Environmental Considerations 13.6.1 System Integration 13.6.2 Solar System Considerations 13.6.3 Solar Collectors 13.6.4 Solar Pond 13.6.5 Photovoltaics 13.6.6 Environmental Impact Nomenclature References

14 Wind Energy Powered Desalination Systems Jaime González, Pedro Cabrera and José A. Carta 14.1 Introduction 14.2 Basic Wind Technology Concepts 14.2.1 Brief Classification of Wind Energy Exploitation Systems 14.2.2 Horizontal-Axis Wind Turbine Components 14.2.2.1 Energy Acquisition Subsystem 14.2.2.2 Mechanical Power Transmission Subsystem 14.2.2.3 Yaw Subsystem 14.2.2.4 Electrical Subsystem 14.2.2.5 Control Subsystem 14.2.2.6 Support Subsystem 14.3 Particular Characteristics of Wind Energy 14.3.1 Wind Resource Estimation 14.4 Classification of Wind-Driven Desalination Systems 14.4.1 On-Grid Wind Energy Systems for Desalination 14.4.1.1 Wind Turbines that Dump all the Generated Energy into the Grid 14.4.1.2 Micro-Grids Interconnected with a Conventional Grid

542 542 545 548 550 550 553 553 553 554 556 556 557 558 559 560 567 568 570 570 574 575 579 583 584 589 589 590 591 598 600 602 606

Contents xvii 14.5 Off-Grid Wind Energy Systems for Desalination 14.5.1 Small-Scale Systems 14.5.1.1 Electrical Interface in the Coupling between Wind Energy and Desalination Unit 14.5.1.2 Mechanical and Hydrostatic Interfaces in the Coupling between Wind Energy and Desalination Unit 14.5.2 Medium- and Large-Scale Systems 14.5.2.1 Electrical Interface in the Coupling between Wind Energy System and Desalination Unit 14.5.2.2 Mechanical and Hydrostatic Interfaces in the Coupling between Wind Energy System and Desalination Unit 14.6 Wind-Diesel Systems for Desalination 14.7 Conclusions and Future Trends List of Symbols References 15 Geothermal Desalination Veera Gnaneswar Gude 15.1 Introduction 15.2 Renewable Energy Powered Desalination 15.3 Geothermal Energy Utilization Around the World 15.4 The Rationale – Why Geothermal Desalination? 15.4.1 Capacity Factor 15.4.2 Comparable Costs 15.4.3 Efficient Resource Utilization 15.4.4 Integrated Uses for Geothermal Energy Sources 15.5 Global Geothermal Desalination Potential 15.5.1 Geothermal Water Composition 15.5.2 Geothermal Water for Thermal Desalination 15.5.3 Geothermal Water for Membrane Desalination 15.6 Geothermal Desalination – State of the Art 15.6.1 Thermal Desalination Processes 15.6.2 Membrane Desalination Processes 15.7 Desalination Process Selection 15.7.1 Plant Size 15.7.2 Geothermal Energy Quality and Quantity and other Renewable Energy Sources

608 608

609

617 620

621

628 630 634 638 639 647 648 649 649 651 652 653 654 655 656 657 659 660 661 661 663 667 667 668

xviii

Contents

15.7.3 Desalination Technology 15.7.4 Feed Water 15.7.5 Product Water 15.7.6 Brine Disposal 15.7.7 Techno-Economic Requirements 15.8 Challenges and Considerations for Geothermal Desalination Implementation 15.8.1 Land Use 15.8.2 Geological Hazards 15.8.3 Waste Heat Releases 15.8.4 Atmospheric Emissions 15.8.5 Water Footprint 15.8.6 Noise and Social Impacts 15.9 Techno-Economics of Geothermal Desalination 15.10 Summary References 16 Future Expectations 16.1 Introduction 16.2 Historical Trends in Fresh Water Supply Development 16.3 Emerging Trends and Directions in Alternative Water Supply Development 16.3.1 Desalination of Impaired Waters 16.3.1.1 El Paso’s Kay Bailey Hutchison Desalting Plant 16.3.2 Impaired Water Usage in Energy Production 16.3.2.1 Palo Verde Nuclear Power Plant, Arizona 16.3.3 Salinization 16.4 Desalination for Oil and Gas 16.4.1 Treatment of Produced Water from Conventional Reservoirs 16.4.2 Designer Waterflooding for Enhanced Oil Recovery 16.4.2.1 Conventional Reservoirs 16.4.2.2 Unconventionals 16.4.3 Treat to Need 16.4.4 Treatment of Hydrofracking Flowback 16.4.5 Water Treatment and the Oil Sands 16.5 The Future of Desalination Technologies 16.5.1 Biomimetic and Nanotech Membranes

668 668 669 669 669 670 671 673 673 673 674 674 674 676 678 683 683 684 687 692 693 696 698 699 700 701 702 702 704 705 706 710 712 715

Contents xix 16.5.2 Desalination with Renewables 16.6 Summary References

716 716 717

List of Contributors

721

Index

737

Preface The world-wide demand for “fresh” water is growing exponentially, while the supply of readily-available fresh water is dwindling. Several diverse techniques have been implemented to try to meet the growing demand for fresh water, with variable degrees of success. One technique that has had great success and that continues to grow in application is desalination. Desalination encompasses a host of technologies such that clean water may be generated regardless of location, make-up source, and/or energy source. This book explores numerous desalination technologies. Some of the technologies that are covered here are highly commercialized and are in extensive use today, while others are under development and may be commercially-viable tomorrow. This book also covers renewable energy sources (wind, geothermal, and solar) as alternatives to fossil-fuels to drive desalination technologies. World-renowned experts have contributed to this book. The authors’ experience includes decades of work in their respective fields, and covers the gamut from academia to real-world practice. I thank the authors for contributing their time and sharing their expertise to help us explore the possibilities within the realm of desalination.

xxi

1 Introduction to Desalination Jane Kucera Nalco Water/an Ecolab Company

Abstract The availability of fresh water on the planet is finite, and natural fresh water makes up only about 0.5% entire water supply on Earth. This limited supply, coupled with the growing population of the Earth and the growing industrialization of many developing countries, is driving global fresh water stress and scarcity to the point where more fresh water must be found to meet future needs. Methods to “find” more fresh water include conservation and reduce/reuse/recycle of existing fresh water sources, moving fresh water from water-rich regions to water-poor regions, and “creating” fresh water from other sources, such as oceans and wastewater, using desalination. Of these methods, desalination has proven to be a very viable technique to meet current and future fresh water needs in many areas around the world. This introductory chapter discusses the history of, and drivers for desalination, and also provides a framework for the detailed discussions about various desalination technologies and opportunities to use renewable energy sources to power the desalination technologies that are presented in this book. Keywords: Desalination, water scarcity, thermal desalination, membrane desalination, reverse osmosis, renewable energy sources

1.1 Introduction Desalination: from the root word desalt meaning to “remove salt from” [1]. By convention, the term desalination is defined as the “process of removing dissolved solids, such as salts and minerals, from water” [2]. Other terms that are sometimes used interchangeably with desalination are desalting Corresponding author: Jane Kucera ([email protected]) Jane Kucera (ed.) Desalination 2nd Edition, (1–49) © 2019 Scrivener Publishing LLC

1

2

Desalination 2nd Edition

and desalinization, although these terms have alternate meanings; desalting is conventionally used to mean removing salt from other more valuable products such as food, pharmaceuticals, and oil, while desalinization is used to mean removing salt from soil, such as by leaching [2]. The first practical use of desalination goes back to the sixteenth and seventeenth centuries, when sailors such as Sir Richard Hawkins reported that their men generated fresh water from seawater using shipboard distillation during their voyages [3]. The early twentieth century saw the first desalination facilities developed on the Island of Curaçao and in the Arabian Peninsula [3]. The research into and application of desalination gained momentum in the mid-twentieth century, and the last 30 years has witnessed exponential growth in the construction of desalination facilities. One could ask the question, “Why desalination?” Desalination has become necessary for several reasons, the most compelling of which may be: 1) the increased demand for fresh water by population growth in arid climates and other geographies with limited access to high-quality, lowsalinity water, and 2) the per capital increase in demand for fresh water due to industrialization and urbanization that out paces availability of highquality water. Research and development over the last 50 years into desalination has resulted in advanced techniques that have made desalination more efficient and cost-effective. Desalination is, and will be in the future, a viable and even necessary technique for generating fresh water from water of relatively low quality. Thus, the title of this book, Desalination: Water from Water. In this chapter, and in this entire book, we make the case for desalination as one of the major tools for meeting the fresh water needs of a growing and industrializing planet.

1.2 How Much Water is There? The allocation of the world’s water is shown in Figure 1.1. About 97.5%, or 1338 million km3, of the world’s water is sea-water [3, 4]. Eighty percent of the remaining water is bound up as snow in permanent glaciers or as permafrost [4]. Hence, only 0.5% of the world’s water is readily available as low-salinity groundwater or in lakes or rivers for “direct” use by humans.

1.2.1 Global Water Availability Some regions of the world are blessed with an abundance of fresh water. This includes areas with relatively low populations and easy access to

Introduction to Desalination 3 Ocean seawater 97.5% Permanent Snow & Ice 80% Ground & surface water 20% Fresh & brackish water 2.5%

Figure 1.1 Allocation of the world’s water resources.

surface waters, such as northern Russia, Scandinavia, central and southern coastal regions of South America, and northern North America (Canada, Alaska) [2, 5]. More populated areas and areas with repaid industrialization are experiencing more water stress, particularly when located in arid regions. There are numerous methods to calculate water stress (e.g., The Faulkenmark Indicator [6]), and many maps that display current and projected future water stress. In most cases, water stress is measured by comparing the amount of water used to that which is readily available, as explained by Maplecroft: “The Maplecroft Water Stress Index evaluates the ratio of total water use (sum of domestic, industrial, and agricultural demand) to renewable water supply, which is the available local runoff (precipitation less evaporation) as delivered through streams, rivers, and shallow groundwater. It does not include access to deep subterranean aquifers of water accumulated over centuries and millennia. The application of the index is to provide a strategic overview of the current situation of physical water stress at global, continental, regional, and national levels. It does not take account [any] future projection, [or] water management policies, such as desalination, or the extent of water re-use” [5]. Figure 1.2 shows the baseline water stress for the world, as estimated by the World Resources Institute for 2015. The areas of the world that are not rich in water resources and that also experience un-stable and rapid population growth and industrialization will

4

Desalination 2nd Edition

Figure 1.2 Global baseline water stress, 2015. Courtesy of World Resources Institute.

see water stress significantly increase in the future. Figure 1.3 compares the global water stress in 1995 with that predicted for 2025 [7]. As many as 2.8 billion people will face water stress or scarcity issues by 2025; by 2050, that number could reach 4 billion people [7] (See Figure 1.4 for world-wide 2040 estimates). Water stressed areas will include the south central United States, Eastern Europe, and Asia, while water scarcity (extremely limited access to flush water) will be experienced in the Southwestern United States; Northern, Southern, and Eastern Africa; the Middle East; and most of Asia [2].

1.2.2 Water Demand The demand for water in developed nations is relatively high. Demand in the United States is about 400 liters per person per day [4]. Some Western countries that have been successful in implementing conservation and reuse measures have seen their demand for water drop to about 150 liters per person per day [4, 8]. However, the limited availability and access to water in some parts of the world, results in much lower consumption in these regions. For example, per capita freshwater consumption in Africa is only about 20 liters per day due to the shortage of suitable water [8]. The

Introduction to Desalination 5

1995

2025 Water withdrawal as a percentage of total available water More than 40 % From 40 % to 20 %

From 20 % to 10 % Less than 10 %

Figure 1.3 Global water stress in 1995 and predicted for 2025. Courtesy of Philippe Rekacewicz (Le Monde diplomatique), February 2006.

Figure 1.4 Projected water stress by 2040. Courtesy of World Resources Institute.

World Heath Organization (WHO) deems 15 to 20 liters per person per day is necessary for sur vival, while 50 liters per person per day is estimated to be needed for operation of basic infrastructure such as hospitals and schools (see Figure 1.5) [4]. The WHO estimates that by 2025, the worldwide demand for fresh water will exceed supply by 56% [8]. In addition to population growth, another pressure being exerted on water supply is fact that the per capita water demand is increasing faster than the rate of population growth [9]. According to Global Water Intelligence [10], the per capital water demand has outpaced population growth by a factor of 2. By 2050, global water demand is expected to increase 55% over

Desalination 2nd Edition Daily per capital demand, liters/person

6

400 350 300 250 Basic infrastructure requirements

200 150

Basic needs

100 50 0 USA

Western nations with conservation

Africa

Figure 1.5 Global demand for water and World Health Organization basic water requirements (2010).[4, 8].

2015 demands, primarly due to manufacturing, thermal electricity generation and domestic use [11].

1.2.3 Additional Water Stress Due to Climate Change While population growth and per capita increase in demand are two major water stressors, the impact of climate change on global water stress cannot be ignored. The effects of climate change actually work synergistically with population growth and increasing demand to strain water supply. As population and industrialization grow, climate change accelerates, leading to more drastic climate events such as drought. A study by the National Center for Atmospheric Research (NCAR) indicates that severe drought is a real possibility for many populous countries [12]. Regions that are projected to experience considerable drought include most of Latin America, the Mediterranean regions, Southeast and Southwest Asia, Africa, the southwest United States, and Australia [9]. Coincidentally, many of these regions are also experiencing increases in population, industrialization and, urbanization, with the corresponding increase in per capita water demand. The United Nations forecasts that the world will have 27 cities with populations greater than 10 million by the year 2020, and all but 3, New York City, Moscow, and Paris, will be in regions under the threat of significant drought [9]. Risks to freshwater supplies increase with increasing greenhouse gas emissions (via industrialization). [11] For example, higher seawater levels due to melting of polar ice can lead to a variety of problems, including

Introduction to Desalination 7 seawater intrusion into coastal aquifers and higher water temperatures, leading to faster dissolved oxygen depletion, both of which affect the quality of this fresh water source. [13]. The effects of climate change on water balance and availability, coupled with population growth and industrialization, will create added future challenges for finding more fresh water to meet demand.

1.3 Finding More Fresh Water For much of the world’s urbanized population, fresh water is an afterthought, a commodity that has been easy to find and always there at the tap. However, water in some parts of the world is increasingly considered a “product” that must to be found and developed to meet growing demand. Depending on the specific circumstances in a particular geography, one or more methods may need to be implemented to find and develop water sources to meet future water needs. Some of these methods are summarized below.

1.3.1 Relocating Water Moving water from water-rich areas to water-scarce regions, while sounding extreme, is not a new idea. Witness the diversion of water to the desert southwest United States for drinking, power, and irrigation uses. Los Angeles currently imports 85% of its water demand from outside sources: the Sierra Nevada Mountains, the Delta in Northern California, the Los Angeles Aqueduct, and the Colorado River Aqueduct [14]. However, moving water is not always palatable. Public outcry against moving water from a water-rich region can be a formidable obstacle. Consider the Columbia River in the Pacific Northwest United States. “Water is Oregon’s Oil,” declared Oregon State Senator David Nelson in his 2007 white paper, “Columbia River Diversion as a Public Revenue Source.” Diversion of the Columbia River to other western states has been a topic of discussion in the State of Oregon for over 40 years. Not much has come of this discussion to date however, as water-poor areas in the region have found other sources for water, and, more to the point, Oregonians have routinely declined to give up their supply of inexpensive fresh water that also serves as their source for relatively inexpensive hydroelectric power. Politics can also play a role in how water supplies are dispersed. In the late 2000’s, different political parties in Spain were having a tug-of-war

8

Desalination 2nd Edition

over how to supply the south-eastern area of Spain with water. The conservative party in Spain advocated moving water from the Ebro River (an eastern river whose delta into the Mediterranean Sea is about half way between Barcelona and Valencia) to the Community of Valencia, which lies approximately 200km from the delta. The Socialist Party in power has commissioned the Torrevieja Seawater Reverse Osmosis (SWRO) facility, the 6th largest SWRO facility in the world, which is located in Alicante, Municipality of Torrevieja, about 75 km from Valencia. Backers of the Ebro river project have denied a permit for concentrate discharge from the SWRO facility, thereby preventing the construction of the seawater intake and outfall pipelines [14]. The Terreveija facility was delayed for 3 years due in part to the political wrangling. Having been finally constructed, the facility is designed to deliver 240,000 m3/ day to approximately 400,000 people (see Figure 1.6). The Valencia province has 2.5+ million people with several more SWRO projects under way that could encounter the same political stalemate. While importing fresh water makes sense in some cases, public and political pressures, as well as technical issues, such as moving water long distances, particularly when elevation changes are involved, will not make importing water supplies feasible or even possible to meet the requirements of all regions in need of fresh water.

Figure 1.6 Terrevieja, Spain, 240,000 m3/day seawater reverse osmosis desalination facility. Courtesy and copyright of Acciona.

Introduction to Desalination 9

1.3.2 Conservation and Reuse Conservation is a term that has been used for decades to mean more efficient usage and savings of a resource, in this case, water. The twenty-first century equivalent terms for conservation are sustainability, and more recently green, and reduce/reuse/recycle. Regardless of which term is used, the need to conserve through more efficient usage, recycling, and reuse has become popular in today’s culture. While these techniques are oft times the first choice of populations located in arid areas or far from an ocean as a means of finding more fresh water, all populations can benefit from these techniques. For example, consider the City of Los Angeles, California, an arid, coastal city that receives only about 40cm of rain a year. Los Angeles imported roughly 85% of its water from northern California, the Owens River, and the Colorado River as of 2013 (see Figure 1.7). Los Angeles is one large metropolitan area that has considered conservation to supply an increasing portion of its future water needs. Los Angeles Country has a current population of about 10.2 million people and is expected to grow to reach 26 million inhabitants by 2060 [16]; water demand is expect to rise by 123 million m3 per year [9, 17]. The Los Angeles Department of Water

Figure 1.7 Water sources for Los Angeles, California, USA [15].

10

Desalination 2nd Edition

and Power (LADWP) describes the future of the city’s water philosophy: “Conservation will continue to be a foundation of LADWP water resource management policy, and will be implemented to the fullest extent concurrent with further consideration of alternative water supplies” [18]. In addition to its aggressive conservation plan, the LADWP has developed a new Recycled Water Master Plan which relies heavily on recycling highly-treated wastewater as a cost-effective solution to meet some of the future demands of the city [19]. The Edward C. Little Water Recycling Facility (ELWRF) located in the City of El Segundo, Los Angeles County, CA (commonly referred to as “West Basin”), is a model for water conservation, recycling, and reuse. The facility, funded in 1992 following the severe drought in California in the late 1980’s and early 1990’s, produced about 236,000 m3/d of recycled water in 2013 at a 2012 investment of $500 million [20]. Five grades of water, known as “designer” water, are produced by the facility to match the needs of local industry (water type listed roughly from lowest to highest quality): 1. Tertiary wastewater (known as Title 22 Water) for general industrial and irrigation uses, such as irrigating golf courses, 2. Nitrified water for use in industrial cooling towers, 3. Softened reverse osmosis (RO) water for ground water recharge, 4. RO water for low-pressure boiler feed water at local refineries, and Table 1.1 Actual and projected water sources for Los Angeles Department of Water and Power. Adapted from [15]. Water source (%) Imported Water, total

LADWP 2010 (actual) LADWP 2035 (plan) 84

64

Metropolitan Water District (MWD)*

48

24

Los Angele Aqueduct

36

34

Other water transfers

0

6

Ground Water

14

16

Conservation

1

9

Recycled Water

1

8

Storm Water Capture

0

3

*Northern California’s Sacramento and San Joaquin Rivers, via the State Water Project, and  the Colorado River, via the Colorado River Aqueduct, provide 45% of MWD water sources.[21]

Introduction to Desalination 11 5. Ultra-pure RO water for high-pressure boiler feed water at local refineries. The objective of the LADWP Recycled Water Master Plan is to recycle a total of about 62 million m3 of water per year by 2019 at an estimated cost of $715 million to $1 billion [14, 18]; by 2035, the goal is to recycle 72 million m3/year [22]. West Basin has already achieved 2/3 of that water recycling goal. The recycled water conserves 42 million m3/year of water that would have to be imported from elsewhere to meet demands [23]. Table 1.1 shows the water sourcing plan for the LADWP from actual sourcing in 2010 to projected sourcing in 2032. [15]. Conservation and recycling wastewater, using West Basin as the example in Southern California, will require treatment, such as desalination, to produce water that is suitable for reuse. Conservation and recycling has the potential to slow the rate at which new, future supplies of fresh water may need to be developed, but will not, by itself, meet the total local and worldwide need for fresh water.

1.3.3 Develop New Sources of Fresh Water Developing new sources of fresh water other than traditional sources, such as lakes, rivers, or relatively shallow wells, is another method for meeting the demand for more fresh water. The most common new sources for developing new fresh water supplies are seawater and deep wells or saline aquifers, and waste water. Seawater is the traditional source water when one thinks of desalination. Seawater represents the feed water source for the majority of desalination facilities in the world (59%) [24]. The majority of these facilities were developed in the Arabian Gulf region, Algeria, Australia, and Spain. In the

Figure 1.8 Claude “Bud” Lewis Carlsbad (California) Desalination Facility. Courtesy of Poseidon Water.

12

Desalination 2nd Edition Saudi Arabia 19%

ROW 29%

47.4 million m3/d total seawater capacity

UAE 18%

India 2% China 3% Qatar 4% Israel 4% Australia 4%

Spain 8% Algeria Kuwait 5% 5%

Figure 1.9 Seawater disalination by country 2015 [25].

EXPLANATION Depth to saline ground water, In feet Less than 500 500 to 1,000 More than 1,000

0

Inadequate information

0 200 400 KILOMETERS

200

400 MILES

Figure 1.10 United States Geological Survey map of depth to saline ground water in United States aquifers, cir. 1965.

united states seawater desalination is also being used to reduce the dependence of the Southern California Region on imported water. Southern California has a handful of direct seawater desalination facilities, the largest of which is the Claude “Bud” Lewis Carlsbad SWRO Desalination Facility near San Diego (see Figure 1.8). The facility was commissioned in December, 2015, and supplies 190,000 m3/day of fresh water to San Diego

Introduction to Desalination 13

Figure 1.11 Observed minimum depth to brackish or highly-saline groundwater in the United States and Selected US Territories for 2017 Survey [USGS]. Courtesy of the U.S. Geological Survey, Professional Paper 1833, https://doi.org/10.3133/pp1833, April, 2017.

County. Another 190,000 m3/day facility is in later stages of installation in Huntington Beach, near Los Angeles. For sea-bounded, arid areas, turning to the sea for water is only natural, (see Figure 1.9 [25]. However, seawater supply is only suitable as a desalination source for coastal areas; inland areas would need to rely on other sources, such as saline aquifers for new water supply. Figure 1.10 shows a United States Geological Survey (USGS) map of US saline aquifers; the map was generated in 1965 and was not updated until 2017 [26]. The 2017 USGS study, Brackish Groundwater in the United States, (Professional Paper 1833, published April, 2017) [27], looked at more than 380,000 sites (compared to 1,000 in the original, 1965 survey). This recent study generated a significantly more comprehensive view of the characteristics of US aquifers than the original survey (see Figure 1.11). The USGS found that groundwater in US aquifers is 800 times the volume of fresh water currently used in the US [26]. For California, desalinated groundwater could provide enough water to meet the state’s needs for 160 years [28]. While some applications don’t require high-quality water, such as mining, oil and gas development, and thermoelectric power generation, and can use the groundwater as is, other applications (including most industrial,

14

Desalination 2nd Edition

Table 1.2 Classification of source waters as a function of total dissolved solids (TDS). Source water

Total dissolved solids (ppm)

Classification

Drinking Water* < 500

Fresh

Fresh

< 1,000

Fresh

Brackish

1,000–5,000

Mildly Brackish

5,000–15,000

Moderately Brackish

15,000–35,000

Heavily Brackish / Seawater

35,000

Standard Average Seawater

35,000–45,000

Seawater

Seawater

* World Health Organization [10].

municipal, microelectronic, and pharmaceutical) would require treatment of the groundwater via desalination prior to use [26]. The key to groundwater use is sustainability; the ability to recharge the aquifers in a reasonable time and manner [28]. Note that most current activity involving saline aquifers centers on using them as storage for greenhouse gases, primarily carbon dioxide, rather than as sources for fresh water [29]. This is presumably due to the need to treat the water to generate fresh water from the saline brines as opposed to the relative ease of injecting greenhouse gases, a process that does not require treatment, into the aquifers. One arid area that is already desalinating aquifers to provide fresh water is the state of Texas in the united states as of March, 2016. Texas had more than 100 desalination facilities, all using brackish ground water as the feed source, although several seawater desalination facilities are planned. [30] While most Texas facilities are small or intermittent operation, the largest facility by far is the Kay Bailey Hutchinson facility in El Paso. This facility can generate approximately 104,000 m3/day. [30] Another and perhaps more challenging feed source is wastewater. Industries and municipalities have wastewater that can be treated and reused for a variety of industrial applications. West Basin in Los Angeles is a good example of using desalination technologies, ion exchange softening and RO, to recover industrial–grade water from wastewater, as discussed previously. Wastewater from various industries, for example, power and refineries, each have their own characteristics, which typically must be treated before discharge. Adding a desalination step to the treatment can, in many cases, yield water suitable for reuse within the facility, thereby reducing the requirement for fresh water from natural sources.

Introduction to Desalination 15 Seawater and other sources present an opportunity to meet the growing water needs of the world. Table 1.2 lists generally-held classifications of water as a function of salinity (note that saline aquifers are generally considered to be at least moderately brackish and most wastewater are at least mildy brackish). Considering these classifications, even the highersalinity “fresh” water would require treatment for potable or industrial use to reduce the concentration of dissolved minerals. Thus, desalination can be used to generate high-quality water from water that is, without treatment, not suitable for direct use.

1.4

Desalination: Water from Water

1.4.1 Drivers for Desalination One can conclude from the discussions in this chapter that new sources of fresh water must be developed to meet the growth in the demand. Apart 120

Contracted

Online Capacity (million m3/d)

100 80

Global cumulative contracted capacity (2016*): 95.6 million m3/d

60

Global cumulative online capacity (2016*): 88.6 million m3/d

40 20 0 1965

1970

1975

1980

1985

Capacity (million m3/d)

* Values through June 2016

1990 1995 Year

2000

2005

2010

2016

8

Contracted - 1st half

7

Contracted - 2nd half

6

Contracted - full year Online

5

Online capacity, 2015 (full year): 3.1 million m3/d

4 3

Contracted capacity, 2015 (full year): 3.2 million m3/d

2

Online capacity, 2016 H1*: 2.1 million m3/d

1 0 1980

1985

1990

* Values through June 2016

1995

2000

2005

2010

2016

Year

Figure 1.12 Growth of cummulative (a) and new (b), on-line desalination capacity. Courtesy of Global Water Intelligence.

16

Desalination 2nd Edition

from moving water from location to location, reuse of wastewater and use of alternate sources of water will require treatment to yield water that is suitable for potable or industrial use. And, since wastewaters and alternate source waters are generally brackish or highly saline, desalination technologies will most certainly be required as part of the treatment scheme. Thus, the driver for desalination is clear: future demand for high quality water will require desalination of water sources that are lower in quality (higher in dissolved solids) than are commonly utilized today (and which may not be available tomorrow). [31] Desalination of various water sources to provide a supply of fresh, usable water has been growing almost exponentially since 1965, when global commissioned desalination capacity was less than 2 million m3/d [32]. By June, 2016, the global commissioned desalination capacity was over 88 million m3/d [33]. Figure 1.12 shows the rate of increase in the cumulative, on-line capacity since 1965. [33] However, the incremental, new, on-line and contracted capacity has waned considerably from the peaks in the late 2000’s (see Figure 1.12). [33] The IDA Desalination Yearbook 2016–2017 explains that the downturn is due, at least in part, from “… low commodity [oil] prices, and the dependence of certain regional economies on these prices, helping to cancel out some of the increasing demand from factors such as population growth, increased industrialization, drought and climate factors, and competition for water resources.” [33]

1.4.2

Feed Water Sources for Desalination

Feed water sources for desalination are varied. As previously discussed, feed sources can range from seawater and saline aquifers to wastewater 6

Seawater Brackish Other

Capacity (million m3/d)

5 4 3

Global cumulative contracted capacity (2016*): 95.6 million m3/d

2

Global cumulative online capacity (2016*): 88.6 million m3/d

1 0 1990

1995

* Values through June 2016

2000

2005

2010

2016

Year

Figure 1.13 Annual new contracted desalination capacity by feed water type [33]. Courtesy of Global Water Intelligence.

Introduction to Desalination 17 Wastewater 6%

Pure Water 5%

River Water 9%

74.8 million m3/d installed capacity Brackish Water 21%

Seawater 59%

Figure 1.14 Total, global installed capacity by feed water source as of 2012 [37]. Courtesy of Global Water Intelligence.

for recycle and reuse. While seawater represents the feed water source for the majority of desalination facilities, the use of other feed water sources, such as brackish water, saline aquifers, and wastewater, has been growing steadily since 2000 [33]. Figure 1.13 shows the growth in annual new contract capacity by feed water type through June, 2016. [33] As the figure shows, the downturn in incremental capacity shown in Figure 1.10 is primarily due to the decrease in demand for seawater desalination plants, particularly in the Middle East and North Africa.[33] Brackish water and other sources showed volatility but continued to show modest increases through June, 2016. Figures 1.14 shows more detail in the total worldwide install capacity by feed water type through August, 2012. [25, 37] Although feed sources for desalination appear to be limited in number, (e.g. seawater, brackish water saline aquifers, and wastewater) composition of specific examples of the various make-up source classifications can differ greatly depending on their hydrologic origin. Table 1.3 demonstrates some of the variability in well and surface waters, with a standard seawater and a sample grey water source included for comparison (the well, river, and grey waters shown in Table 1.3 either are currently being used as feed water sources for desalination facilities or have been considered for use as source water for such facilities) [3]. Despite variations in quality among the various feed water sources, they all share the characteristic of being relatively high in salinity or total dissolved

10,752

2,701

1,295

416

390

145

66

8

6.4

4.5

3.0

1.3

Sodium

Sulfate

Magnesium

Calcium

Potassium

Bicarbonate

Bromide

Strontium

Silica

Boron

Nitrate

Fluoride

0.0034

19,345

Chloride

Iron

35,000

Standard seawatera

Total Dissolved Solids

Species (ppm)

NA

0.61

0.11

NA

29.4

NA

0.05

75

15.9

176

38.4

301

745

1,370

3,170

Well waterb

0.01

NA

1.7

NA

29

1.6

NA

220

2.8

50

16

110

38

3.9

350

Well waterc

NA

0.5

2.6

NA

14.2

1.4

NA

160

5.7

104

38.3

342

185

181

1,021

Colorado river waterd

NA

1.1

1.7

NA

2.6

0.13

NA

134

1.7

39

12

27

8.3

16

210

Lake michigan watere

Table 1.3 Sample water composition of seawater, well water, surface water, and grey water sources.

1.3

0.56

4.4

NA

42

1.2

NA

244

21

88

34

200

66

67

650

Grey waterf

18 Desalination 2nd Edition

0.003

0.0004

0.00009

Uranium

Manganese

Selenium

e

NA

NA

NA

NA

0 0023

NA

0.0038

0.0035

NA

NA

NA

NA

Water from Chicago, Illinois, Anne Arza, Nalco Water, an Ecolab Company, personal communication, May 16, 2011

Colorado River Water near Andrade, CO [3]

Hotel hygiene wastewater, Las Vegas, Nevada, August 21, 2007

f

NA

NA

NA

NA

Water from Benson, Arizona, Kevin O’Leary, Aquatech, personal communication, April 7, 2008

d

c

El Paso Water Utilitites, El Paso Airport Wells [3]

References 3, 4, and 8

b

a

NA = Not available

0.003

Arsenic

0.03

NA

NA

Introduction to Desalination 19

20

Desalination 2nd Edition

Other 1% Military 1% Tourism 2% Irrigation 2% Power 6%

Municipal 60%

85.2 million m3/d Total nationwide installed capacity

Industry 28%

Figure 1.15 Total, 2015 global installed capacity by type of user [35]. Courtesy of Global Water Intelligence.

Other 0.5% Hybrid 0.8%

CEDI 0.3%

NF/SR 2%

ED 3.6%

ED/EDR 3%

MED 8.0%

MSF 26.8%

MED 7%

65.2 million m3/d installed capacity

MSF 21% RO 60%

(a)

Other 2%

86.5 million m3/d total worldwide installed capacity

RO 65%

(b)

Figure 1.16 Global installed desalination capacity by technology for 2010(a) and 2016(b) [24, 32]. Courtesy of Global Water Intelligence.

solids (TDS). High salinity (and, in some cases, other mineral constituents) makes the water unsuitable for direct potable and industrial use. Therefore, demineralization or “desalination” treatment to reduce the concentration of TDS must be part of the treatment system employed if these sources are to be used to supplement or replace existing fresh water supplies.

1.4.3 Current Users of Desalinated Water The primary user for desalinated water is the municipal sector; sixty percent of desalinated water is used for potable application. [35] Industrial and

Introduction to Desalination 21 power users together accept another third of the worldwide desalination capacity (see Figure 1.15) [35]. Although potable applications account for nearly twice the total volume of desalinated water used than industrial applications, the number of industrial facilities (including power) out numbered municipal facilities by almost 2 to 1 (8,715 to 4,415 respectively) in 2011, indicating that the size of industrial desalination facilities are considerably smaller than municipal facilities [18]. The remaining 6% of desalinated water is used for irrigation, tourism, military, and other applications [35].

1.4.4 Overview of Desalination Technologies The world-wide installed desalination capacity in 2016 was about 86.5 million m3/d [24]. Figure 1.16 shows the relative breakdown of installed capacity of various desalination technologies for 2010(a) and 2016(b) [24, 32]. Membranes are currently outpacing traditional thermal technologies in total installed desalination capacity. Prior to 1980, membrane technologies made up less than a third of the global desalination capacity [19], while today, membranes account for just roughly 2/3 of the total installed desalination capacity [24]. As shown in Figure 1.16, installed capacity of RO grew from 60% of total installed capacity in 2010 to 65% in 2015, an increase of 8% over the year, while installed capacity of thermal, multi-stage flash (MSF) evaporation decreased by over 21% for the same period; installed capacity of thermal, multi-effect distillation (MED) declined slightly from 8% to 7% remained steady at 8%. Membrane technologies such as RO offer the advantage of smaller infrastructure, and RO total treated water cost is becoming competitive with traditional thermal processes [36].

Capacity (million m3/d)

100

Membrane Thermal

80 60

Installed membrane capacity, 32014: 59.0 million m/d

40 20 0 1980

Installed thermal capacity 32014: 24.0 million m/d 1985

1990

1995

2000

2005

2010

2015

Figure 1.17 Growth of installed membrane and thermal desalination capacity. Courtesy of Global Water Intelligence.

22

Desalination 2nd Edition

Figure 1.18 The Ras Al Khair desalination and power generating facility in Saudi Arabia. The desalination facility generates 1.025 million m3/day of fresh water using MSF and SWRO.

Figure 1.19 The 636,000 m3/day Jebel Ali MSF desalination plant and power complex in Dubai, UAE. Expansion of the desalination plant will use SWRO.

Membrane-based systems are popular in rising markets such as Algeria, Spain, and Australia, while thermal processes are found in traditional markets such as the Middle East, where energy costs are lower [32]. The proportion membrane-based desalination has been growing, as shown in Figure 1.17; globally, new membrane capacity was 93% of total new desalination capacity from 2015 through June, 2016. [33]. Top countries by total installed thermal capacity from 1945–2012 include (approximate, in million m3/d): [37]

Introduction to Desalination 23 United Arab Emirates (UAE): 7.3 Saudi Arabia: 5.9 Kuwait: 2.3 Qatar: 1.3 Libya: 0.9 Top countries by total installed membrane capacity from 1945–2012 include (approximate, in million m3/d): [37] United States: 7.9 Saudi Arabia: 5.1 Spain: 4.6 China: 2.0 UAE: 1.5 The global trend toward membrane systems, primarily SWRO, over traditional thermal processes is driven by a several factors, including [33]: Lower RO capital costs due to less-expensive construction materials; Versatility in RO feed sources (MSF and MED are impractical for all but sweater desalination); No thermal energy requirement for RO, while MSF and MED desalination are cost-effective in areas with subsidized energy or in conjunction with an industrial process, such as power generation that yields inexpensive, low-pressure steam. Growth of membrane technology over thermal processes has occurred even in the traditional thermal markets. In the years 2000–2009, the portion of new membrane capacity in the Middle East and North Africa (MENA) was 55% as compared to 45% for thermal capacity; during 2010– 2014, new membrane capacity increased to 76% of new, installed capacity; and in in the period of 2015 through June, 2016, new, membrane capacity had increased to 86% of the total new capacity. [33] While thermal plants still predominate in the Middle East, due to easy access to fossil fuel resources (relatively inexpensive energy), the relatively poor seawater quality (high salinity (up to 55,000 ppm TDS)) in the Gulf Region that limits RO recovery, and high the temperature of the seawater, which promotes biofouling of the membranes, has made the water in this region not well suited for previous generations of SWRO [38]). However,

24

Desalination 2nd Edition

as pretreatment and membrane materials have improved (e.g., higher permeability and salt-rejecting membranes that operate at lower pressure), SWRO has become viable in the Middle East. Energy-recovery devices now used with SWRO also reduce the mechanical energy requirements of the process [39, 58]. Further, energy initiatives in the region, such as the Dubai Clean Energy Strategy 2050, currently drive new desalination capacity in the region toward SWRO [41]. New membrane capacity in the region is typically in the form of hybrid MSF/SWRO facilities. The largest desalination facility in the world, Ras Al Khair Desalination Plant, in Saudi Arabia employs both MSF and RO to generate 1.025 million m3/day of water (the facility is co-located with a 2,650 MW combined cycle power plant—see Figure 1.18). The 8 MSF units and 17 RO units generate 727,000 m3/day and 309,000 m3/day, respectively, of fresh water. The Jebel Ali desalination facility in Dubai, UAE, which also co-located with a power station (see Figure 1.19) currently generates 636,000 m3/day of fresh water using 8 USF units, but is looking to expand capacity by 182,000 m3/day using SWRO. Dubai Electricity and Water Authority (DEWA) chose RO for the expansion due to its lower-energy requirements. [41] Indeed, many thermal plants in the region are expanding capacity using SWRO.

1.4.5 History of Desalination Technologies Desalination has grown substantially since the mid 1960s. In 1952, there were only about 225 desalination facilities world-wide with a total capacity of about 100,000 m3/d. [2] In 2015 there were over 15,000 desalination facilities globally [32] and in 2016, the total global capacity was over 86 million m3/d; 44% of this capacity was located in the Middle East and North Africa [24, 31]. While there are many desalination technologies in use or being developed today, desalination began using thermal processes. Membrane-based processes, such as RO, helped to further promote desalination over the last 50 years. The history of these pioneering technologies is outlined below.

1.4.5.1

History of Thermal Desalination

Thermal desalination techniques were recognized as early as 320 B.C. when Aristotle wrote, ‘saltwater, when it turns into vapor, becomes sweet and the vapor does not form saltwater again when it condenses.’ Shipboard distillation beginning in the sixteenth century is the first practical use of distillation to generate fresh water from seawater [42]. In 1843, Rillieux successfully patented, built, and sold multi-effect evaporators [42].

Introduction to Desalination 25 The number of thermal desalination installations has grown rapidly over the last 100 years. However, while RO and other membrane technologies were revolutionary in development, the development of thermal desalination technologies over the last 45 years has been more evolutionary than revolutionary [43]. Multi-effect distillation was the first thermal desalination technology employed [42, 43]. The first units were designed with submerged tube evaporators the exhibited low heat transfer rates and high scaling rates. Vertical- and horizontal-tube evaporators (also known as falling-film evaporators) used in modern MED facilities provide higher heat transfer coefficients and lower specific heat transfer surface area requirements than their older counterparts. Drawbacks of the current MED technology are the complexity and production capacities [43]. Also, the relatively low, maximum-brine temperature of MED (~65 oC) due to scale-forming issues, is another limitation of MED. However, the use of membrane pretreatment such as nanofiltration (NF) prior to MED to remove the calcium that contributes to calcium-sulfate scale in MED units has been considered as a way of allowing higher temperature operation of MED and, thereby enhance the use of MED for desalination [43]. Due to the early issues with MED (e.g., scaling and low heat transfer rates), MSF distillation technology was developed in the late 1950s and early 1960s as an alternative. Flashing distillation was first commercially employed by Westinghouse in Kuwait in 1957 [42]. That same year, an

Figure 1.20 MSF Desalination plant (110,000 m3/day capacity) combined with power plant - Doha West, Kuwait. Courtesy of MEW (Ministry of Electricity and Water).

26

Desalination 2nd Edition

MSF distillation patent was issued; in 1959/60, the first commercial MSF facilities were installed in Kuwait (19 stages, 4550 m3/d) and the Channel Islands (40 stages, 2775 m3/d) [44, 45]. In 1973, “standard” MSF units, that produce justover 27,000 m3/day and consisted of 24 stages, were developed [42]. In 1984, the Doha West Power and Desalination plant came on line in Kuwait City, Kuwait, with a desalination capacity of 110,000 m3/day using 4 MSF units (27,500 m3/day each), see Figure 1.20. This currently, the desalination plant is being augmented with 140,000 m3/day of SWRO capacity (see discussion of hybrid desalination facilities in Section 1.4.4.). Recent developments in the thermal desalination technology have focused on scale and corrosion control techniques and on the increase in distiller production capacity [43]. Early, pre 1980, MSF units were primarily constructed using carbon steel for the shell and the internals [46]. Corrosion of the metal due to seawater resulted in the use of thicker materials of construction, which made the units larger and heavier. Units built after 1980 use stainless and duplex stainless steel to reduce corrosion, allowing for lighter and smaller MSF units. Future advances in the technology will most likely focus on improvements in thermodynamics and material selection [46].

1.4.5.2

History of Reverse Osmosis Desalination

While the earliest recognition of thermal desalination was a few hundred years B.C., the earliest recorded documentation of semi-permeable membranes was in 1748, when the phenomenon of osmosis was observed by Jean-Antoine “Abbe” Nollet [47]. Osmotic phenomenon was also studied in the 1850’s, and then in the 1940’s, when Dr. Gerald Hassler began investigation of the osmotic properties of cellophane [48]. “Modern” RO technology truly began in the late 1950’s when C.E. Reid and E.J. Breton at the University of Florida and Sidney Loeb and Srinivasa Sourirajan (working under prof. Samuel Yuster) at the University of California at Los Angeles (UCLA) independently demonstrated RO using polymeric membranes. Figures 1.21 and 1.22 show Sidney Loeb and the “Big Dripper” flat-sheet, cellulose acetate membrane he developed. The United States was the early leader in desalination research in the 1960s and 70s due in most part to government funding. The Saline Water Conversion Act of 1952 established the Office of Saline Water (OSW) in 1955, which later became the Office of Water Research and Technology (OWRT) in 1974. It was under such governmental programs that Loeb and Sourirajan developed the first commercially-viable RO membrane while at the UCLA [48]. In 1965, the tubular, cellulose acetate membrane

Introduction to Desalination 27 developed at UCLA became the membrane used in the first commercial RO facility located in Coalinga, California [49]. Figure 1.22 shows UCLA team at the Coalinga facility. Government funding also lead to the development in 1965 of the solution/diffusion membrane transport model by Harry Lonsdale, U. Merten, and Robert Riley at the General Atomic Division of General Dynamics, Corp [50]. This model has become the basis of research and development of new membrane materials since that time. [51] It was under similarly-funded governmental programs that John Cadotte, while at North Star Research, prepared the first interfacial polyamide membrane US patent US 454,039,440, issued August 7, 1977, and later while at Filmtec corporation, the US patent US 4,277,344, issued July 7, 1981, and commonly referred to as the “Cadotle 344” patent, that soon after became the basis of the FilmTec FT30 membrane (now part of DowDupont) [52]. The original FT30 membrane chemistry is the basis of the majority of reverse osmosis membranes in use today [44]. The OWRT was abolished in 1982 and government funding of desalination research in the United States dropped considerably. By that time, however, much of the foundation for RO, membrane-based desalination had been laid. Since then, incremental membrane improvements have been made in the areas of flux, rejection, and operating pressure requirements, as shown in Table 1.14 [45]. However, no major “breakthroughs” in terms of higher membrane selectivity with higher water flux and chlorine tolerance have occurred in 45+ years since the revolutionary early developments. Research is continuing, however, and the development of nanotechnology and “nanocomposite” membranes circ. 2005 has raised hopes that RO membranes with higher selectivity, water flux, and chlorine tolerance may be on the horizon, all of which would reduce costs and improve efficiency associated with membrane desalination [44].

1.4.5.3 Developments in Desalination Since 1980 Since 1980, the world-wide development of desalination techniques has been driven out of necessity due to water scarcity and population growth. The private sector has led the investment in research and development as they began to see water not as a commodity, but as a product to be sold at a profit [3]. This development by the private-sector has lead to a significant drop in cost of water generated through desalination techniques. An example of such is the 80% reduction in price of RO membrane elements over the last 30 years, while incremental improvements in flux, selectivity, and operating pressure were realized (Tables 1.4 and 1.5 [43, 52]). In 1991, the

28

Desalination 2nd Edition

Figure 1.21 Sidney Loeb (a and b) and coworker (b) with Loeb’s “Big Dripper” flat sheet membrane machine. Courtesy of Julius Glater.

Figure 1.22 UCLA Team at Rain Tree facility, Coalinga, California, USA. Courtesy of Julius Glater.

cost to produce water at the SWRO Charles E. Meyer Santa Barbara desalination facility was about 2.00/m3; in 2007, the estimated cost had dropped to about 0.89/m3 [2, 46, 56] (The facility was developed in the wake of the severe drought of the late 1980’s, and operated from March to June, 1992. However, sufficient water supplies since 1991, caused the facility to go into long-term, standby mode. The plant was reactivated in 2015, again due to severe drought conditions, and became operational in May, 2017. The

Introduction to Desalination 29 Table 1.4 Advances in brackish water reverse osmosis membrane performance [45]. Year

Pressure (bar)

Relative flux

Rejection (%)

1970’s

30

1

97

1980’s

20

1.9

99.0

Cross-linked polyamide composite

1987

15

3.0

99.7

Cross-linked aromatic polyamide composite

1988

10

4.2

99.7

Cross-linked aromatic polyamide composite

1996

7.6

5.6

99.7

Cross-linked aromatic polyamide composite

1999

5.2

8.0

99.7

Cross-linked aromatic polyamide composite

2017*

5.2

8.0

99.8

Cross-linked aromatic polyamide composite

Membrane material Cellulose acetate

* Hydranautic CPA7

Table 1.5 Decline in membrane cost relative to 1980 [55]. Year

Relative cost

1980

1.00

1985

0.65

1990

0.34

1995

0.19

2000

0.14

reactivated facility will produce 11.4 m3/day using 40% less energy than the original design, due to retrofitting the facility with more efficient pumps/ motors and membranes. [53]. Figure 1.23 shows the general cost range of desalinated water for MSF(a) and RO(b) (variability in cost is due to factors such as size of plant, degree of pretreat-ment employed and makeup source in the case of RO) [34, 39]. By comparison, direct municipal

30

Desalination 2nd Edition 12.0

Unit cost ($/m3)

9.0

6.0

3.0

0.0 1955

1960

1965

1970

1975

1980

1985

1990

1995

2000

2005

Year 6.0

5.0

3

Unit cost ($/m )

4.0

3.0

2.0

1.0

0.0 1965

1970

1975

1980

1985

1990

1995

2000

2005

Year

Figure 1.23 Cost distribution reduction as a function of time for MSF (a) and RO (b) processes. RO process costs vary with source water TDS, and generally show higher costs for seawater and wastewater, followed by river water and brackish water at the lower cost range [34]. Courtesy of Elsevier.

water from easily accessable wells and surface sources cost about $0.40/m3 in 2012 [54]. There have been other desalination techniques invarious stages of development, some commercially successful, but none more commercially successful than traditional thermal and RO desalination processes. Table 1.6 lists a selection of desalination technologies and their current status.

Commercial application limited to kidney dialysis Mature

Osmotic pressure

Electric Field

Osmotic pressure

Temperature (vapor pressure)

Dialysis

Electrodialysis / Electrodialysis Reversal (ED / EDR)

Forward Osmosis (FO)

Membrane Distillation (MD)

On-going study of fundamentals; little industrial-scale study to date [60]

Gaining commercial application for desalination, particularly since 2010 [58]

Mature

Electric Field

Continuous Electrodeionization (CEDI)

Status

Membrane

Driving force

Technology

Technique

(Continued)

Capital cost of capillary membranes used for MD is higher than RO membranes, thereby limiting commercial application [57, 61]

Requires additional treatment such as low-pressure RO to yield purified water suitable for use. Differs from PressureRetarded Osmosis (PRO) in that PRO Is used primarily to generate power. [59]

Best suited to charged, low molecular weight species; less efficient with low salinity feed streams or when low-conductivity product streams are required.

Osmotic process is inherently not selective and transport by diffusion is slow such that application to desalination Is not economically feasible. [57]

Used primarily for polishing RO permeate; some process (non-water purification) applications.

Comment

Table 1.6 Sample desalination technologies. Technologies covered in this volume are noted in italics.

Introduction to Desalination 31

Temperature

Temperature

Evaporation

Pressure

Reverse Osmosis (RO)

Carrier-gas

Pressure

Piezodialysis

Thermal

Pressure

Nanofiltration (NF)

Membrane

Driving force

Technology

Technique

Table 1.6 Cont.

Mature

Under Development

Mature

Little commercial interest today [57]

Mature

Status

Surface scaling, improved heat transfer, and corrosion- resistant materials primary focus of current development [46]

Also known as humidification dehumidification desalination (HDH); process has high specific energy requirements, usually related to the dehumidification step [62]

Research into fouling resistance and chlorine tolerance and well as reduction in energy while maintaining or improving rejection is in progress.

Researched in the 1970s [59]. Salt separation (enrichment) decreases with increasing feed salt concentration; uses charge mosaic (ion exchange) membranes.

Used for “softening” or partial demineralization of salt water and as pretreatment to desalination using RO.

Comment

32 Desalination 2nd Edition

Other

Electric field

Ionic Charge

Ion Exchange

Temperature

Capacitive Deionization (CDT)

Freezing-Melting (FM)

Mature

Under Development since the 1970s

Under Development, currently limited to food processing where high-temperature treatment could adversely affect product quality

Limited to relatively low-salinity brackish water desalination applications.

Older carbon aerogel electrodes (made from resorcinol) were expensive and their ion storage capacity was relatively low (64); newer hierarchical carbon aerogel monoliths show improved performance [65]

Relatively high capital cost and process complexity; potential for desalination using freezing-melting technique in concert with other desalination technologies [63]

Introduction to Desalination 33

34

Desalination 2nd Edition

1.4.6

The Future of Desalination

Desalination today is still a capital and energy-intensive proposition. Methods to reduce costs are necessary to make the desalinated water more affordable. To that end, developments to increase the efficiency (and reduce costs of desalination) are needed. (Note that while some predictions call for a reduction of desalinated water costs of over 50% within 20 years [31], others predict that, while operating efficiencies will improve, little, if any, decrease in the costs for desalinated water will be realized [35]; instead, an increase in conventional water treatment and water importation costs is projected [35], which may help offset the cost of desalination.) As of 2014, there were at least 5 national research initiatives underway, and over 50 active universities with desalination-related research & development (R&D) programs, plus hundreds of private sector R&D project ongoing [35]. Some areas of current development include: Energy: There has been considerable reduction in the energy required to drive SWRO desalination plants, (Figure 1.24), due to improvements in membrane permeability and energyefficient pumps, and the use of energy-recovery devices. [40]. In 2008, a pilot-scale SWRO system demonstrated desalination at 1.8 kWh/m3 [66] using high-permeability seawater elements. Work by the Desalination Coalition in 2006 [67, 68] demonstrated seawater desalination of 1.58 kWh/m3 under very ideal conditions, e.g., new membranes, no fouling, low flux, and 42% recovery.[69] These two examples include energy only for desalination, excluding energy for plant intake, pretreatment, post treatment, and brine discharge. Current SWRO plants in operation consume about 2.5–3.5 kWh/m3 [70]. In 2016, the Perth, Australia SWRO plant used 3.5 kWh/m3 of produced water (including total energy, from intake to customer), using wind energy and advanced energy recovery systems. [70] However, it is projected that the likelihood of any major, future improvements in in organic energy efficiency of SWRO is small [35]. The theoretical energy required to recover 50% of 35,000 ppm TDS seawater (by any desalination method) is 1.06 kWh/m3, and assumes a reversible thermodynamic process. [40]. In actuality, energy consumption will be higher since desalination plants do not operate as a reversible thermodynamic process [40].

Power consumption (KWh/M3)

Introduction to Desalination 35 20 –

4–



3–

15 –

2–



1–

10 –

0– 2000

2004

2008

– 5– – 0– 1970

1980

1990

2000

2004

2008

Year

Figure 1.24 The change in power consumption for RO in SWRO plants from the 1970s to 2008. The horizontal dashed line corresponds to the theoretical minimum energy required for desalination of 35,000 ppm seawater at 50% recovery (1.06 kWh/m3). The energy data exclude the energy required for intake, pretreatment, posttreatment, and brine discharge. [65]

Renewable energy to drive desalination projects, such as wind, solar, and geothrmal of are being considered to reduce the energy footprint also of desalination (today’s contribution of renewable energy sources to desalination processes is only about 0.05%) [71]. Table 1.7 shows some of the renewable energy sources (RES) currently considered to power desalination processes. Table 1.8 lists key data for RES and CO2-free desalination [72]. Materials: materials of construction for thermal processes that resist corrosion and reduce the size and weight (and costs) of units need to be developed. New membrane materials that resist attack by chlorine and fouling with microbes and that also show higher selectivity and higher flux are projected to be developed [31, 35]. Higher premeability will lead to lower energy costs for RO operations as discussed above. Although the development of new membrane materials, including nanocomposites, carbon nanotubes, aquaporins,

Limited to relatively small, less than 100 m3/h installations; most significantly smaller. [74]

First Geothermal desalination plant was constructed in 1972, in the US. Fairly mature technology [77]

Heat can be used directly in thermal desalination or indirectly via electrical production to power RO. [77]

Under Development for desalination since 1982 [74]

Wind Energy–RES

Used to generate electricity (using techniques known as “Delbuoy” RO concept, “McCabe” wave pump or “water hammer”) which can then be used to power RO desalination. [74]

Glothermal Energy–RES

Under Development since 1980s; relatively few desalination installations today [74]

Wave Energy–RES

Can be used to power thermal desalination process directly and to generate power for RO desalination processes.

Can reduce CO2 emissions generated by thermal desalination plants by about 90% [76] compared to those using fossil fuels

One of the first desalination power sources used to generate fresh water (via evaporation of salt water) [73]

Solar Energy–RES

Comments

Nuclear Energy–CO2 free source Proven for power generation; growing application for cogeneration of electricity and desalinated water [75]

Status

Renewable energy source

Table 1.7 Renewable energy sources (RES) and CO2-free technologies used to replace traditional fossil-fuel based energy sources to drive traditional desalination processes. These play a relatively small role in desalination today [71], but given the concerns regarding fossil fuels (e.g., availability, carbon footprint (see Table 1.9)), these sources show potential for future application in powering desalination. Solar, wind and geothermal energies are covered in this volume.

36 Desalination 2nd Edition

2.6–6.5

Production Cost: $/m3

‡‡‡

seawater

brackish water

electrodialysis

‡‡



**photovoltaic

*membrane distillation

100

1–100

1.5

Electrical Energy: kWh/m3

Typical Capacity: m3/d

Application, R&D

Development Status

Thermal Energy: kJ/Kg

Solar MED

Item

10.4–19.5

0.1–10

< 200

0

R&D

Solar MD*

Table 1.8 Key data for renewable energy desalination, 2013 [72].

6.5–9.1 BW 11.7–15.6 SW

< 100

0

0.5–1.5 BW 4.0–5.0 SW‡‡‡

‡‡

Application, R&D

PV**/RO

10.4–11.7

< 100

0

3.0–4.0 BW

R&D

PV/ED‡

3.9–6.5 BW 6.5–9.1 SW

50–2000

0

0.5–1.5 BW 4.0–5.0 SW

Application, R&D

Wind RO

Introduction to Desalination 37

38

Desalination 2nd Edition and graphene, will continue, it is projected that these will not replace conventional, thin-film composite membranes for all but niche applications (e.g., high-purity or polishing). [35]. Chemicals: antiscalants for both thermal and membranebased desalination processes to increase the degree of water recovery per given unit size, while reducing the potential for scaling. Optimization: as energy costs will no doubt increase in the future, it is important that desalination plants are efficiently designed and operated as close as possible to the limits of performance of components, such as the RO membranes. [35]

Despite thise development areas, no major breakthroughs are expected over the several years to radically lower the cost of desalination; instead, improvements in technologies will lead to slow but steady improvements. [31, 55] Current desalination techniques can also have a significant impact on the environment, as described below [71]. Before desalination can become sustainable, these environmental issues must be addressed. The two most pressing issues are concentrate disposal and airborne emissions: Concentrate disposal—very high salinity wastewater is generated through thermal and RO membrane desalination. Total dissolved solids can be as high 100,000 ppm in the wastewater from the desalination of seawater. Furthermore, concentrate can be highly turbid and be at elevated temperatures (thermal desalination plants), and can contain chemical additives such as polymers/coagulants, acid, biocides, corrosion inhibitors, and cleaners. The issue becomes how to dispose of the wastewater in an environmentally-safe manner that is also cost effective. Seawater desalination facilities currently discharge to the ocean, which some argue is damaging to local flora and fauna via increases in sea-water temperature, salinity, and turbidity. These conditions may be harmful to marine life and cause them to migrate away, and, at the same time, enhancing the populations of algae and nematodes.[78, 79] Inland brackish water facilities use other disposal methods, including discharge to surface water (45%), discharge to sewer (27%), deep well injection (16%), land application (8%), and discharge to evaporation ponds (4%) [79]. Each of these disposal methods has

Introduction to Desalination 39 its own environmental concerns that are directly related to the high salinity and other characteristics of the water to be discharged. Carbon footprint and airborne emissions—Table 1.9 lists airborne emissions for fossil fuel-powered desalination technologies. The carbon dioxide emissions from thermal desalination processes are a full order of magnitude higher than that for RO, when powered by fossil fuels. And, the energy to power the desalination facilities has the largest impact on the carbon footprint of the process [80]. The emissions for desalination technologies can be reduced significantly when these processes are powered by waste heat or RES such as solar radiation rather than by fossil fuels as shown in Tables 1.8 and 1.10 [71]. Future demand for fresh water has led to a total desalination market value of 12.8 billion in 2015 [25] and will likely result in an estimated $19.9 billion expenditure on global desalination projects in the year 2020 [25]. Figure 1.25 shows the actual and projected market overview from 2000 to 2030. [25] Total, installed capacity in June, 2015 was 86.8 billion m3/d; project installed capacity for 2030 is over $50 billion. It is apparent that RO will be the primary mode of desalination for the foreseeable future, [33], Figure 1.14 [24] demonstrates the relative growth of RO to traditional thermal processes. Thermal processes, while in decline, still have a foothold, primarily in the Gulf Region Particularly where cogeneration with power yields low-cost steam. [24]

Table 1.9 Cir. 2010 airborne emissions per cubic meter of water generated by various desalination technologies when powered by fossil fuels and/or waste heat [71]. MSF Emissions

MED

RO

Fossil Fuel Waste Heat Fossil Fuel Waste Heat Fossil Fuel

CO2 (kg)

24

2.0

18

1.1

1.8

Dust (g)

2.0

2.0

1.0

1.0

2.0

NOx (g)

28

4.1

21

2.4

3.9

NMVOC (g)

7.9

1.2

5.9

0.6

1.1

SOx (g)

28

15

26

16

11

MED

RO

10

4.2

20

NOx (g)

NMVOC (g)

SOx (g)

*100 kWp (peak power output) NA = not applicable

11

8.8–11.9

0.5–0.85

1.2–2.5

0.7–1.1

6

0.34

0.8

0.3–0.4

20

3

7

8.2

15

0.2–0.4

0.9–1.5

0.4–0.7

NA

NA

12

NA

0.15 NA

0.7

0.3

5–8

0.4–0.7

1.2–2.1

0.5–0.9

2.1

0.2

0.4

0.2

Solar Solar Solar Solar Solar Solar thermal photovoltaic* Wind thermal photovoltaic* Wind thermal photovoltaic* Wind

CO2 (kg)

Emissions

MSF

Table 1.10 Emissions for desalination processes powered by RES [82].

40 Desalination 2nd Edition

3,000

60,000

2,500

50,000

2,000

40,000

1,500

30,000

1,000

20,000

500

10,000

0 2000

2005

2010

* Values through June 2016 Cumulative contracted

2015 Year

2020

Annual Contracted

2025

Cumulative capacity (MGD)

Annual capacity (MGD)

Introduction to Desalination 41

0 2030

28th Inventory

Figure 1.25 Annual and cumulative global desalination capacity as a function of time as calculated and projected by Global Water Intelligence and DesalData.com. Courtesy of Tom Pankratz.

While desalination plays a vital role in the sustenance and growth in many arid and water-poor areas of the world, there are still barriers to more widespread implementation, particularly for seawater desalination. For desalination to be a sustainable method to develop sources of fresh water in the future, desalination technologies need to be economically and energy efficient and have a low environmental impact. Indeed, concerted efforts on these fronts have resulted in significant advances in desalination [81], which are described in this book: improvements in performance and design of traditional technologies, development of innovative new technologies, the marriage of desalination technologies with RES.

1.5

Desalination: Water from Water Outline

In this volume we present the case for desalination, describe conventional and innovative new desalination technologies, present RES options for desalination, and conclude with a discussion of future directions. In addition to conventional desalination technologies discussed in Section I (thermal and RO), there are many other technologies under development, as listed in Table 1.6. Current and several of the more

42

Desalination 2nd Edition

promising desalination technologies under development are discussed in Sections II through IV of this book: Section II: this section covers traditional thermal desalination technologies, including MED, and MSF. These technologies are very mature, but do have limitations that may be overcome with through future development of new materials to improve corrosion resistance and heat transfer, and through the development of antiscalants. Section III: this section describes several membrane-based technologies, including RO, continuous electrodeionization, and membrane distillation, as well as some membrane-based desalination technologies that have only recently been commercialized, namely forward osmosis. Significant research is on-going to surmount the limitations of relatively mature membrane technologies such as RO, while also developing less commercialized technologies such as membrane distillation. Section IV: this section details non-traditional desalination technologies. Technologies covered in this section include freezing-melting desalination processes, capacitive deionization, and ion exchange. These technologies may be limited to niche desalination applications [35], and some require additional development to become viable for widescale commercial or industrial desalination. The future need for RES to power desalination may be as great as the need for desalination itself, if the issues associated with fossil fuels become even more acute. Figure 1.26 shows possible corolations between RES and desalination technologies [83–85]. Solar, wind, and geothermal energies are described in Section V as alternatives to fossil fuels to power desalination. The book concludes with a discussion about the future prospects for desalination in Section VI. This final section discusses future water sources for desalination, including traditional sea-water sources, and other, more impaired sources, such as industrial wastewater. Future water demand for desalination water, including traditional municipal users and emerging water users, such as oil field hydrofracking, is profiled. Finally, research needs to develop additional desalination technologies that are efficient and cost effective are presented along with some of the more promising desalination techniques to come out of research and development.

Introduction to Desalination 43

Figure 1.26 Relationship between renewable energy source and disalination technologies.

Abbreviations BWRO

brackish water reverse osmosis

CDT

capacitive deionization technology

CEDI

continuous electrodeionization

ED

electrodialysis

EDR

electrodialysis reversal

ELWRF FM FO HDH kWp LADWP MD MED MENA

Edward C. Little Water Recycling Facility freezing-melting forward osmosis humidification dehumidification peak power output, kilowatts Los Angeles Department of Water and Power membrane distillation multi-effect distillation Middle East and North Africa

44

Desalination 2nd Edition

MEW MSF MWD NCAR NF OSW

Ministry of Electricity and Water (Kuwait) multi-stage flash Metropolitan Water District of Southern California National Center for Atmospheric Research nanofiltration Office of Saline Water

OWRT PRO PV RES RO SWRO TDS UAE UCLA

Office of Water Research and Technology pressure-retarded osmosis photovoltaic renewable energy sources reverse osmosis seawater reverse osmosis total dissolved solids United Arab Emerites University of California at Los Angeles

USGS WHO

United States Geological Survey World Health Organization

References 1. Merriam Webster Dictionary http://www.Merriam-Webster.com. 2001. Accessed October 12, 2012. 2. National Research Council, Committee to Review the Desalination and Water Purification Technology Roadmap, Review of the Desalination and Water Purification Technology Roadmap, The National Academies Press, Washington, D.C., www.nap.edu/catalog/10912. html, 2004. 3. National Research Council, Committee on Advancing Desalination Technology, Desalination, A National Perspective, The National Academies Press, Washington, D.C., 2008. 4. Micale, G. A. Cipollina, and L. Rizzui, Seawater Deslination for Freshwater Production, Seawater Desalination, Conventional and Renewable Energy Processes, G. Micale, A. Cipollina, and L. Rizzui (Ed.), Springer, Heidelberg, Germany, (2009). 5. McGeown, J. Personal Communication, March 5, (2012). 6. Faulkenmark, M., “The Massive Water Scarcity Threatening Africa—Why Isn’t It Being Addressed,” Ambio 18, no. 2 (1989) 112–118. 7. Rekacewicz, Philippe, http://www.grida.no/grphicslib/detail/ increasedglobal-water-stress_5694, accessed October 15, 2012.

Introduction to Desalination 45 8. The World Health Organization, Water Facts & Water Stories from Across the Globe, http://www.theworldwater.org/water_thefacts.php (2010). 9. Kerschner, M. Edward, and M W. Peterson, Peak Water: The Preeminent 21st Century Commodity Story, Morgan Stanley Smith Barney Global Investment Committee, November, (2011). 10. Water Desalination Report. Global Water Intelligence, March, (2010). 11. United Nations, “Water for a Sustainable World,” United Nations World Water Development Report 2015, (2015) http://unesdoc.unesco.org/ images/0023/002318/231823E.pdf, accessed March 31, 2017. 12. Dai, Aiguo, Drought Under Global warming: A Review, Wiley Interdisciplinary Reviews: Climate Change 2, October, (2010). 13. IPCC (Intergovernmental Panel on Climate Change), “Climate Change 2014: Impacts, Adaptation, and Vulnerability,” Working Group II Contribution to the Fifth Assessment Report of the Intergovernmental Panel on Climate Change, Cambridge University Press, Cambridge/New York, (2014). 14. SWRO Falls Victim to Politics, World Desalination Report, vol. 47, number 19, 16 May, (2011). 15. Faith2Green, “LADWP Energy Efficiency and Water Conservation Effort,” http://www.faith2green.org/ladwp-energy-efficiency-and-water-conservation-effort.html, accessed March 16, 2017. 16. Cox, Wendell, “California in 2060?” New Geography, May 22, 2015, http:// www.newgeography.com/content/004926-california-2060, accessed march 16, 2017. 17. LADWP, Securing L.A.’s Water Supply, “City of Los Angeles Water Supply Action Plan, http://www.ladwp.com/ladwp/cms/ladwp010587.pdf, May, (2008). 18. LADWP, http://www.ladwp.com/ladwp/cms/ladwp001620.jsp, (2011). 19. Mabry, Kat, City Seeks Public Input on New Recycled Water Program, Encino Patch, http://www.encino.patch.com/articles/city-seeks-public-input-onnew-recycled-water-program, May 19, (2011). 20. West Basin Municipal Water District http://www.westbasin.org/ water-reliability-2020/recycled-water/water-recycling-facility, accessed October 15, (2012). 21. Metropolitan Water District of Southern California (MWD) Website, “Sources of Water,” http://www.mwdh2o.com/AboutYourWater/Sources%20 Of%20Supply/Pages/default.aspx, accessed March 16, 2017. 22. Los Angeles Department of Water and Power (LADWP) Website, “Recycled Water,” https://www.ladwp.com/ladwp/faces/ladwp/aboutus/a-water/a-wrecycledwater?_adf.ctrl-state=yiso9mqq4_4&_afrLoop=678434215814290, accessed March 16, 2017. 23. Parsons website, “Edward C. Little Water Recycling Facility - Phase V Expansion,” https://www.parsons.com/projects/Pages/eclwrf.aspx, accessed March 16, 2017. 24. Pankratz, Tom, IDA Desalination Yearbook, 2015–2016, for Global Water Intelligence, Media analytics, Ltd. Publishers, Oxford, United Kingdom, (2015).

46

Desalination 2nd Edition

25. Pankratz, Tom, Global Water Intelligence, personal communication March 13, 2017. 26. USGS (United Stage Geological Survey), “USGS Assessment of Brackish Water Could Help Nation Stretch Limited Freshwater Supplies,” April 5, 2017, https://www.usgs.gov/news/usgs-assessment-brackish-water-could-helpnation-stretch-limited-freshwater-supplies, accessed May 22, 2017. 27. Stanton, J.S., D.W. Anning, C.J. Brown, R.B. Moore, V.L. McGuire, S.L. Qi, A.C. Harris, K.F. Dennehy, P.B. McMahon, J.R. Degnan, and J.K Bohlke, Brackish Groundwater in the United States, USGS Professional paper 1833, https://doi.org/10.3133/pp1833. 28. Jaffer, Naizam, “USGS Finds Vast Reserves of Salty Water Underground in California,” The Water Network, April 29, 2017, https://thewaternetwork.com/ article-FfV/usgs-finds-vast-reserves-of-salty-water-underground-in-california-rkPApPg70RUYuHX4gHdNQA, accessed May 22, 2017. 29. Mission 2013: Carbon Sequestration MIT, http://igutek.scripts.mit. edu/terrascope/index.php?page=index, accessed October 15, (2012). 30. Texas Desalination Association Website, “Common Questions about Water Desalination,” http://www.texasdesal.com/about-us/desal-faqs/, accessed March 17, 2017. 31. Voutchkov, Nikolay, “Desalination—Past, Present, and Future,” International Water Association, August 17, 2016, http://www.iwa-network.org/desalination-past-present-future/, accessed March 10, 2017. 32. Pankratz, Tom, IDA Desalination Yearbook 2010-2011, for Global Water Intelligence, Media Analytics, Ltd. Publishers, Oxford, United Kingdom, (2010). 33. Pankratz, Tom, IDA Desalination Yearbook, 2016–2017, for Global Water Intelligence, Media analytics, Ltd. Publishers, Oxford, United Kingdom, (2016). 34. Zhou, Yuan, and Richard S.J. Tol, “Implications of Desalination to Water Resources in China—an Economic Perspective,” Desalination 164 (3) (2004), p225–240. 35. Pankratz, Tom, IDA Desalination Yearbook, 2014–2015, for Global Water Intelligence, Media analytics, Ltd. Publishers, Oxford, United Kingdom, (2014). 36. Pankratz, Tom and T. John , deslination.com, an environmental primer, 2nd ed., Lone Oak Publishing, Houston, Texas, (2004). 37. Pankratz, Tom, IDA Desalination Yearbook 2011-2012, for Global Water Intelligence, Media Analytics, Ltd. Publishers, Oxford, United Kingdom, (2011). 38. Greenlee, Lauren F. , Desmond F. Lawler, Benny D. Freeman, and Benoit Marrot,” Reverse Osmosis Desalination: Water Sources, Technology, and Today’s Challenges,” Water Research 43 (2009), 2317–2348. 39. R.V. Reddy & N. Ghaffour, “Overview of the Cost of Desalinated Water and Costing Methodologies,” Desalination 205, (2007) p 340–353.

Introduction to Desalination 47 40. Elimelech, M., and W. A. Phillip, “The Future of Seawater Desalination: Energy, Technology, and the Environment,” Science 333, P. 712–717, August 5, 2011. 41. Water.Desalination+Reuse, “Dubai Prepares to add Desalination Capacity at Jebel Ali Power Station,” https://www.desalination.biz/news/0/Dubaiprepares-to-add-desalination-capacity-at-Jebel-Ali-power-station/8515/, accessed March 20, 2017. 42. Al-Mutaz, S. Ibrehim, Water Desalination, http://faculty.ksu.edu.sa/ Almutaz/ Documents/ChE-413/Desalination_Introduction.pdf. 43. Hamed, A. Osman, Evolutionary Developments of Thermal Desalination Plants in the Arab Gulf Region, http://www.swcc.gov.sa/files/assets/ R e s e arch / Te ch n i c a l % 2 0 Pap e rs / T h e r m a l / EVOLU T IONA RY % 2 0 %20DEVELOPMENTS%20%20OF%20%20%20THERMAL%20 DESALINATION%20%20PLANT.pdf, Beruit Conference, (2004). 44. Kucera, Jane, Reverse Osmosis, Industrial Applications and Processes, Scrivener Publishing, Salem, Mass., (2010). 45. Norman Li, Anthony Fane, W.S. Winston Ho, and M. Takeshi , Advanced Membrane Technology and Applications, Eds. John Wiley & Sons, Inc., Hoboken, NJ, (2008). 46. Sommariva, Corrado, and R. Vekates, MSF Desalination: A Review of its Development, Watermark, issue 18, December, (2002). 47. Cheryan, Munir, Ultrafiltration and Microfiltration Handbook, 2nd ed. 1998 48. Koenigsberg, Diana, Water Warriors, UCLA Magazine, http://www. magazine.ucla.edu/features/waterwarriors, July 1, (2006). 49. Glater, Julius, Professor Emeritus, UCLA, personal communication, February 24, (2009). 50. H.K. Lonsdale, U. Merten, and R.L. Riley, Transport properties of cellulose acetate osmotic membranes. J. Appl. Polym. Sci 9, (1965). 52. Cadotte, John, R.S. King, R.J. Majerle, and R.J. Peterson, Interfacial synthesis in preparation of reverse osmosis membranes. J. Macromol. Sci. Chem 15, (1981). 53. City of Santa Barbara Website, “Desalination,” http://www.santabarbaraca. gov/gov/depts/pw/resources/system/sources/desalination.asp, accessed March 16, 2017. 54. Water Reuse Association, “Seawater Desalination Costs,” White Paper, revised January, 2012. https://watereuse.org/wp-content/uploads/2015/10/ WateReuse_Desal_Cost_White_Paper.pdf, accessed March 17, 2017. 55. Baker, Richard, Membrane Technology and Applications, 2nd ed., John Wiley & Sons, Ltd., Chichester, West Sussex, England. (2004). 56. Desalination, County of Santa Barbara Public Works Online. http:// www. countyofsb.org/pwd/water/desalination.htm, January 27, (2007). 57. Baker, Richard, Membrane Technology and Applications, 3nd ed., John Wiley & Sons, Ltd., Chichester, West Sussex, England. (2012). 58. Thompson, A. Neil. and P.G. Nicoll, Forward Osmosis Desalination: A Commercial Reality, IDAWX/PER11-198, presented at the 2011 IDA World Congress, Perth, Australia, September. (2011).

48

Desalination 2nd Edition

59. F.B. Leitz and W.A. McRae, An engineering analysis of piezodialysis, Desalination 10, issue 3, (1972). 60. Alkhudhiri, Abdullah, D. Haif, and H. Nidal, Membrane Distillation: A Comprehensive Review, Desalination 287, (2012). 61. Gullinkala, Tilak, D. Brett, G. Colleen, H. Richard and C.E. Isabel, Desalination: Reverse Osmosis and Membrane Distillation, in Sustainable Water for the Future, E. Isabel and S. Andrea (Ed.), Elsevier, Amsterdam, The Netherlands, (2010). 62. Narayan, G. Prakash, H. R H. McGovern and H.J. Lienhard, Helium as a Carrier Gas in Humidification Dehumidification Desalination Systems,” ASME Conference Proceedings, Paper number IMECE2011-62875, http:// dx.doi.org/10.1115/IMECE2011-62875, accessed November 28 (2012). 64. Dietz, Steven, Improved Electrodes for Capacitive Deionization, SBIR contract DMI-0216299, presented at Proceedings of the 2004 NSF Design, Service and Manufacturing Grantees and Research Conference, Birmingham, AL, January (2004). 65. http://www.rwlwater.com/capacitive-deionization-promising-desalinationtechnology-resdesigned/, accessed November 28 (2012). 66. MacHarg, J., T. F. Seacord, and B. Sessions, “ADC Baseline Tests Reveal Trends in Membrane Performance,” Desalination and Water Reuse 18, p. 30–39, 2008. 67. Veerapaneni, S., B. Long, S. Freeman, and R. Bond, “Reducing Energy Consumption for Seawater Desalination,” J. Amer. Water Works Assn 99, P. 95–106, 2007. 68. Cooley, Heather, Peter Gleick, and Gary Wolff, “Desalination, With a Grain of Salt—A California Perspective, June, 2006, http://pacinst.org/publication/desalination-with-a-grain-of-salt-a-california-perspective-2/, accessed March 17, 2017. 69. Shannon, M. A., P. W. Bohn, M. Elimelech, J. G. Georgiadis, B. J. marinas, and A. M. Mayes, “Science and Technology for Water Purification in the Coming Decades,” Nature 452, P. 301–310, March 20, 2008. DOI: 10.1038/ nature06599. 70. American Membrane Technology Association (AMTA), “Membrane Desalination Power Usage Put in Perspective,” White Paper (FS-7), April, 2016. 71. Al-Ansari, S. Mohammed and Nader Al-Masri, Future Sustainable Desalination Technologies for the GCC, Presented at the Third SQU-JCCP Joint Symposium, Muscat, December. (2010). 72. International Renewable Energy Agency (IRENA), “Water Desalination using Renewable Energy,” Technology Policy Brief I12, January, 2013. 73. M.T. Chaibi and Ali M. El-Nashar, Solar thermal Processes, A review of solar thermal energy technologies for water desalination, in Seawater Desalination, A. Cipollina, G. Micale and L. Rizzuti (Ed.), Springer-Verlag, Berlin, GE. (2009).

Introduction to Desalination 49 74. E. Tzen, Wind and wave energy for reverse osmosis, in Seawater Desalination, A. Cipollina, G. Micale and L. Rizzuti (Ed.), Springer-Verlag, Berlin, GE. (2009). 75. World Nuclear Library, updated March, 2017, http://www.world-nuclear. org/information-library/non-power-nuclear-applications/industry/nucleardesalination.aspx, accessed March 31, 2017. 76. B.M. Misra and I. Khamis, Nuclear desalination, A review of desalination plants coupled to nuclear power stations, in Seawater Desalination, A. Cipollina, G. Micale, and L. Rizzuti (Ed.), Springer-Verlag, Berlin, GE, (2009). 77. Azevedo, Francisco Diogo Abreu Santos Moniz, “Renewable Energy Powered Desalination Systems: Technologies and Market Analysis,” Master’s Thesis, Universidade De Lisboa, 2014, http://repositorio.ul.pt/bitstream/10451/15792/1/ulfc112531_tm_Francisco_Azevedo.pdf, accessed March 16, 2017. 78. Birkett, Jim, Desalination at a Glance, IDA White Paper, http:// www.idadesal. org. (2011). 79. Mickley, Mike, Concentrate Management, in The Guidebook to Membrane Desalination Technology, Mark Wilf (Ed.), Balaban Desalination Publications, L’Aquila, Italy (2007). 80. Emerging Trends in Desalination: A Review, UNESCO Centre for Membrane Science and Technology, University of New South Wales, Waterlines Report Series No. 9, October. (2008). 81. http://www.desalination.biz/news/news_story.asp?id=6318&title= Patents+ indicate+’significant+advances’+in+desalination%2C+says+ report, posted January 31, 2012, accessed December 11, (2012). 82. Raluy, R.G., Luis M. Serra, and Javier Uche, “Life Cycle Assessment of Desalination Technologies Integrated with Renewable Energies,” Desalination, DOI: 10.1016/j.desal.2005.04.023, November, 2005. 83. Shatat, Mahmoud, Saffa Riffat, and Said Ghabayen, “State of Water Desalination Technologies using Conventional and Sustainable Energy Sources,” Presented at the 4th International Engineering Conference Towards Engineering of the 21st Century, 2012. 84. Eltawil, Mohamed, Zhao Zhengming, and Liqiang Yuan, “Renewable Energy Powered Desalination Systems: Technologies and Economics—State of the Art,” Presented at the Twelfth International Water Technology Conference, 2008. 85. Eltawil, Mohamed, Zhao Zhengming, and Liqiang Yuan, “A Review of Renewable Energy Technologies Integreated with Desalination Systems,” Renewable and Sustainable Energy Reviews 13 (2009), p 2245–2262.

2 Thermal Desalination Processes Joachim Gebel Department of Technology and Bionics at the Hochschule Rhein-Waal, Kleve, Germany

Abstract This chapter starts with an extract of fundamentals of engineering science such as thermodynamics and heat transfer. Building on this, mass- and energy balances for single-effect and multiple-effect distillation processes are introduced. A complete set of design equations for MED, MSF and mechanically as well as thermally driven vapor compression plants is presented. In order to be able to compare the different processes in terms of energy demand, the so-called Gained Output Ratio as a performance indicator is introduced and discussed. Based on a vivid description of history of thermal seawater desalination, future prospects and challenges for thermal desalination technologies are discussed at the end of the chapter. Keywords: Mass- and Energy Balances, Single-Stage Evaporation, Multiple-Effect Distillation (MED), Multi-Stage-Flash - Evaporation (MSF), Multiple-Effect Distillation with Thermally Driven Vapour Compression (TVC), Single-Stage Evaporation with Mechanically Driven Vapour Compression (MVC), Gained Output Ratio (GOR), Performance Ratio, Primary Energy Consumption, Historical Review

2.1 Introduction Desalination of seawater or brackish water is achieved by several desalination techniques out of which we are going to discuss desalination using thermal energy, i.e. heat, within the frame of this chapter. First, the following conventional thermal-based desalination techniques are discussed in detail: Multiple-Effect Evaporation, also called Multiple-Effect Distillation

Jane Kucera (ed.) Desalination 2nd Edition, (51–138) © 2019 Scrivener Publishing LLC

51

52

Desalination 2nd Edition

or usually abbreviated as MED, Multi-Stage-Flash Evaporation (MSF), MED with thermally driven vapour compressor (MED-TVC), and SingleStage Evaporation with mechanically driven vapour compressor (MVC). Recent developments in thermal desalination and future prospects especially regarding the trend towards hybrid plants are presented in the second part of the chapter. All theoretical considerations are strongly based on the First and Second Law of Thermodynamics. This approach is very reasonable, because it enables the reader to transfer the presented solutions to more complicated problems. The processes are analysed using an elaborated mass- and energy balance as well as the determination of the thermal efficiency. The old-fashioned and most often used Gained Output Ratio (GOR) is by our opinion not an appropriate scale to compare the different technologies to desalt seawater by using thermal energy. Since thermal energy needs to be specified not only by amount but also by pressure level and temperature level, a more sophisticated approach is needed. Our approach uses the primary energy demand, or more thermodynamically spoken, the exergy of the thermal energy.

2.2 2.2.1

Mass- and Energy Balances Single-Stage Evaporation

As a result of the large difference in the vapour pressures, a single-stage process is sufficient to separate water and salt. Theoretically the salt concentration in the distillate produced by evaporation is zero; however, in practice it is always higher than zero as a result of the seawater droplets entrained. As the salt content in the distillate in comparison with the feed concentration in the seawater can be kept low using a well-operating demister, in the following mass balance the salt concentration has been set to zero. It must be emphasised at this point that this is only valid for thermal processes, in reverse osmosis we may always assume that there will be some flow of salt through the membrane, so that it is not possible to set the permeate concentration to zero. Figure 2.1 shows a schematic of a single-stage evaporation process. There are three mass flows entering and leaving the system, which are subscripted as feed F, brine B, and distillate D. If it is assumed that the relevant pumps are outside of the system there are only two energy fluxes to be found: heat supply for the production of the vapour (subscript “H”) and heat extraction in the condensation of the vapour (subscript “C”). In a first

Thermal Desalination Processes

53

QC

mD xD TD Vapour

QH mB xB TB mF xF TF

Figure 2.1 Schematic of a single-stage evaporation.

step we would like to make a mass balance, a salt balance and an energy balance for the system using the following assumptions: 1. Pure water is evaporated. 2. The temperature of the feed flow is the same as the temperature in the evaporation chamber, i.e. the feed is already preheated. 3. There is no subcooling of the distillate or of the brine. 4. The specific heat capacity of feed, brine and distillate is constant and equal to the specific heat capacity of pure water at the respective temperature. With these assumptions we obtain the following equations: 1. Mass balance

mF mB mD

(2.1)

m F x F mB x B

(2.2)

2. Salt balance

54

Desalination 2nd Edition 3. Energy balance as given by the 1st law of thermodynamics

QH

QC

(2.3)

The maximum salt content in the brine is determined by the solubility of the salts. In order to avoid salt precipitation leading to encrustations and blockages (keyword: scaling), the concentration in the brine must be limited. The concentration factor CF being defined as the ratio of brine concentration to feed concentration is usually used as the parameter which gives the limitation of the evaporation, i.e. concentration process:

CF

mF

xB xF

(2.4)

mB

It may immediately be derived from the salt balance (Eq. 2.2) that the concentration factor can also be interpreted as the ratio of the feed mass flow to the brine mass flow. As the capacity of a seawater desalination plant is defined by the distillate production, it is helpful for our analysis to introduce the yield or recovery rate φ. Starting with the definition of the concentration factor CF (Eq. 2.4) we obtain:

CF

mF mB

mF mF mD

1 m 1 D mF

1 1

CF 1 CF

(2.5)

(2.6)

The significance of the concentration factor and the yield may be illustrated by the following example. Assumed figures are: xF = 35,000 mg/l(Mediterranean Sea) (Max. concentration in brine due to scaling) xB = 60,000 mg/l xF = 35,000 mg/l

xB = 60,000 mg/l

CF

xB xF

60, 000 35, 000

1. 7

CF 1 CF

1. 7 1 1. 7

0. 4

This means: 40% of the seawater, which is pumped into the plant, pretreated and heated up to evaporation temperature, will flow out of the plant

Thermal Desalination Processes

55

as a product, in other words as distillate. 60% are rejected as brine. For a plant in the Arabian Gulf the situation is different: xF = 42,000 mg/l xB = 60,000 mg/l

(Arab Gulf) (Max. concentration due to scaling)

CF

xB xF

60, 000 42, 000

CF 1 CF

1.43 1 1.43

1.43 0.3

Here only 30% of the seawater can be recovered as distillate, 70% are rejected as brine. This means extra cost for each m3 of product due to additional pumping energy and additional chemicals for the pre-treatment. In conclusion it may be said that: The concentration factor and the yield should be chosen to be as high as possible. The limiting factor is the salt content, which leads to precipitation, i.e. scaling. A significant increase in both of these parameters can only be achieved in the end by avoiding scaling, for example by the use of chemicals or through the selective removal of the salts, which form the salt-crusts (for example by using Nanofiltration). The energy balance for the system represented above gives evidence that the heat which is introduced will also be released in the condenser. This is in line with the First Law of Thermodynamics, which postulates the conservation of energy. However, for the production of fresh water from seawater the amount of energy supplied is of most particular interest. If only the subsystem “vapour production” is observed, the following balance can be made:

QH

mD

hV ,TV

(2.7)

In this case, Δhv is the heat of evaporation for pure water under evaporation pressure or at evaporating temperature. If evaporation takes place under ambient pressure, then the specific heat demand of single-stage evaporation may be derived as:

QH mD

hV ,100

C

2, 257

kJ kg

627

kWh to

Now, in order to gradually approach the reality of seawater desalination it is necessary to distance ourselves from the following assumption: 1. It is not pure water, which is evaporated but rather salty seawater.

Desalination 2nd Edition

56

2. The feed flow has ambient temperature and has to be preheated to evaporation temperature. The schematic of this extended plant is shown in Figure 2.2. While nothing changes in the mass and salt balance using these assumptions, the heat flow for the preheating of the feed to evaporation temperature now appears in the energy balance. In taking this into account it must QC

mD TD Vapour

TV

QEvap

QH

mB TB

QBPE

TBPE

TPH

QPH

mSW = mF TSW

Figure 2.2 Schematic of a single-stage evaporation process with seawater preheating.

Thermal Desalination Processes

57

be noted that the evaporation temperature in the seawater is increased as a result of the saline nature in comparison with pure water. In terms of figures the preheating of the feed flow may be divided into the preheating to boiling point of the pure water (given by the evaporation pressure in the evaporation chamber) and the additional heat input as a result of the elevation of the boiling point.

QH QH

mD QH

QEvap QPH

(2.8)

mF cP (TV TSW ) mF cP

hV ,TV mD

QBPE

hV ,TV

mF cP ( TPH

TBPE )

TBPE

(2.9) (2.10)

If the concentration factor CF is introduced as defined above, the specific heat demand of a single-stage seawater desalination plant with feed preheating can be expressed as:

QH mD

hV ,TV

CF c ( TPH CF 1 P

TBPE )

(2.11)

According to the 1st rule of thermodynamics the heat which is introduced has to be emitted. This happens on the one hand through the brine having an increased temperature equivalent to the elevation in boiling point. On the other hand the vapour is overheated by the elevation in boiling point. The heat being emitted is therefore composed of evaporation heat and the superheating. In a multiple-effect plant the superheating is reduced by losses on the path to the next stage, i.e. the vapour condenses in the evaporator of the next stage as saturated vapour at the relevant evaporation temperature. Hence the overall heat balance exactly reads as follows:

QH

QC QLosses

How high the thermal energy demand of such a plant is may be illustrated in the following example. Given are the following figures: Seawater temperature: TSW = 20 °C Evaporation temperature: TV = 100 °C Heat of evaporation: Boiling point elevation: Concentration factor:

hV

2, 257

ΔTBPE = 1.0 K CF = 1.4

kJ kg

58

Desalination 2nd Edition Specific heat capacity:

cP

4. 0

kJ kgK

(seawater )

These figures lead to:

QH mD

2, 257

QH mD

kJ kg

2, 257

kJ 1. 4 ((100 20) 1.0) K 4. 0 kgK 1. 4 1

kJ kJ 1,134 kg kg

3, 391

kJ kg

942

kWh to

For a plant with a capacity of, for instance, 15,000 to/d, a thermal energy demand of almost 600 MW may thus be computed!

QH

mD

QH mD

15, 000

to 1 d kWh 942 d 24 h to

588, 750kW

Without going into the details of the costs of fuel required to provide this amount of heat, it may be said that the production of fresh water in such a seawater desalination plant would be extremely expensive and thus uneconomic. Thus, additional measures are required to reduce the specific heat demand: 1. Reduction in heat demand of preheating by recovery of condensation heat 2. Reduction of heat demand by reducing of the evaporation temperature (evaporation under vacuum) 3. Reduction of heat demand for the evaporation using multiple effects To 1: Recovery of condensation heat In Figure 2.3 the flow scheme of a single-stage plant with energy recovery is shown. The seawater flows through the condenser and absorbs the latent heat of the vapour. A part of this water leaves the system and flows back into the sea as “reject” (mRej), the rest is introduced into the plant as feed (or so-called “make-up”). The heat balance for the condenser is as follows:

mD

hV ,TV

mSW cP (Tout ,Cond TSW ) (mRe j mF ) cP

TPH ,Cond

(2.12)

Thermal Desalination Processes

Tout, Cond

mRej

59

TSW mSW

mF QPH TTTD Heating steam

mD TD Vapour

TV

TV

QH THS

QEvap TV+ TBPE

QBPE

mF

Figure 2.3 Flow diagram of a single-stage evaporation process with energy recovery.

As the amount of distillate, i.e. the product of the plant, should normally be maintained at a constant figure, the reject flow may be adjusted according to the inlet temperature of the seawater (summer or winter operation). The feed flow is coupled to the capacity of the plant through the concentration factor CF and is thus given. The reject flow related to the distillate volume is determined as:

hV ,TV

mRe j mD

cP

TPH ,Cond

CF CF 1

(2.13)

The temperature at the condenser outlet is below the condensation temperature of the vapour as a result of the temperature difference required for the heat transfer according to the Second Law of Thermodynamics. Figure 2.4 illustrates this process by use of a temperature vs. heat transfer area diagram. The temperature difference at the so-called hot end of the heat exchanger is called Terminal Temperature Difference ΔTTTD. ΔTTTD is a figure, which should be stated in the design phase from both an economic and technical point of view. As may be seen from the chart, the condenser area increases with a decreasing difference in temperature. Values

60

Desalination 2nd Edition THS TV

TV

TV

Temperature

ΔTTTD Tout, Cond

TSW 100 %

100 %

Area preheater

0

Area condenser

Figure 2.4 Temperature vs. area diagram for the condenser and the preheater.

for ΔTTTD between 2 and 6 Degree K are appropriate for modern seawater desalination plants. In order to preheat the feed to evaporation temperature an additional external heat flux is needed to close the gap between the temperature at the condenser outlet and the evaporation temperature in the evaporation chamber (see Figure 2.4). Eq. 2.10 and Eq. 2.11 may thus be modified as follows:

QH

mD

mF cP ( TTTD

TBPE )

(2.14)

CF c ( TTTD CF 1 P

TBPE )

(2.15)

hV ,TV

and

QH mD

hV ,TV

If a terminal temperature difference of 3 Degree K is assumed, the heat demand in a single-stage evaporation plant thus reduces using the values listed above by more than 30% according to the following calculation:

QH mD QH mD

2257

2257

kJ kg

kJ 1. 4 (3 1.0) K 4. 0 kgK 1. 4 - 1

kJ kJ 56 kg kg

2313

kJ kg

642.5

kWh to

Thermal Desalination Processes

61

Despite this reduction, the amount of heat that amounts to about 650 kWh per tonne of fresh water is too high in order to be able to utilise the process commercially. A remedy for this is the use of the multipleeffect distillation principle.

2.2.2 Multiple-Effect Evaporation Before going into the design of Multiple-Effect-Distillation (MED) plant, it should be clearly stated once again that the multiple-effect nature of a seawater desalination plant should not be confused with the multiple-stage nature of a distillation column. In the latter case the multiple-stage nature is necessary in order to be able to generate a pure product, as the vapour pressures of the elements to be separated are close to each other (e.g. water – alcohol). In seawater desalination a pure product, i.e. salt-free water, can be produced in one stage – the multiple-effect nature is solely necessary in order to minimise the energy demand. Figure 2.5 shows the simplified scheme of a multiple-effect evaporation plant. In order to explain the principle of the multiple-effect nature more clearly no feed preheating has been shown, i.e. it is assumed that the feed flow, which enters the first stage, has already reached evaporation temperature. It is similarly assumed that the distillate produced in the individual stages will be collected and drained off without any heat recovery. The concentrate flows to the subsequent stage without taking into regard the flashing, which will be explained later. In a Multiple-Effect-Distillation plant the succeeding stage serves as condenser for the steam generated in the previous stage. Thereby the first stage is supplied with energy from external sources, normally with heating steam from a boiler or waste steam from a steam turbine. The amount of distillate produced in a plant with N stages is the sum of the distillates of the individual stages, i.e.

mD

mD ,1 mD ,2 ... mD ,N 1 mD ,N

(2.16)

Disregarding flashing, superheating and losses, the amount of heat transferred in each condenser/evaporator unit simply is the product of distillate mass flow and heat of evaporation:

Q1

mD ,1

hV ,TV ,1

Q2

mD ,2

hV ,TV ,2 (2.17)

.... QN QN

1

mD , N mD , N

1

hV ,TV , N

hV ,TV , N

1

Desalination 2nd Edition

mF Q1 QEvap

TV,1

1

mB,1

mD,1

Q2

TV,2

2

mB,2

mD,2

Q3

TV,3

3

mB,3

mD,3

QN

TV,N

N

mB,N

mD,N

mF

mSW

mRej

62

Figure 2.5 Simplified process flow diagram of a multiple-effect distillation plant (MED). (Feed already preheated, without flashing).

Thermal Desalination Processes

63

Assuming energy conservation in all stages, the heat fluxes are equal, so that the following equation is valid:

Q1

Q2

QN 1

QN

QEvap

(2.18)

The specific heat demand of a multiple-effect plant may can be expressed thus with the help of the Eqs. 2.16 and 2.17 as follows:

QEvap

QEvap

mD

mD ,1 mD ,2 ... mD ,N 1 mD ,N

QEvap

QEvap Q1 hV ,TV ,1

mD

Q2 hV ,TV ,2

QN 1 hV ,TV , N 1

QN hV ,TV , N

1

1

hV ,TV , N 1

hV ,TV , N

...

With Eq. 2.18 it follows that:

QEvap

1 1

1

hV ,TV ,1

hV ,TV ,2

mD

...

This equation may be further simplified if an average value for the evaporation heat is introduced. The following equation for the specific energy demand of a Multiple-Effect-Distillation plant may then be valid:

QEvap

hV ,TV ,m

mD

N

with

TV ,m

TV ,i

(2.19)

N

Assuming that the average temperature in a multiple-effect evaporation plant is 40 °C (the temperature profile of the plant will be examined later), and that the number of stages N equals 10, a specific energy demand may be calculated as:

QEvap mD

hV ,TV ,m N

2, 407 10

kJ kg

240.7

kJ kg

66.86

kWh to

The specific thermal energy demand of a Multiple-Effect-Distillation plant, as shown in Figure 2.6, is a hyperbolic function of the number of

64

Desalination 2nd Edition

Specific heat consumption [kJ/kg]

3,000 2,500 2,000 1,500 1,000 500 0 0

3

6 9 Number of stages [N]

12

15

Figure 2.6 Specific heat consumption of a Multiple-Effect-Distillation plant.

stages, whereby the greatest percentage savings may be made in the step from one to two stages. At first glance it appears worthwhile, when examining the process from the energy-saving point of view, to select the highest possible number of stages. The constraint is given similarly to comparable energy-saving measures through the investment cost, which are needed. (The insulation of a tube should not be so thick that the cost for the insulation material and its fitting consume the cost savings on the energy side.) As every extra stage in a multiple-effect plant increases the investment costs approximately linear, this results in the optimisation task represented in Figure 2.7,which says: At which number of stages is the sum of investment costs and energy costs at a minimum? In order to calculate the investment cost of an MED plant it is necessary to know the size of the heat transfer areas of the evaporator/condenser unit. Here we would like to move on to discussing the following questions: 1. How is the preheating of the feed flow to be incorporated into the multiple-effect plant? 2. How can the heat, which is contained in the distillate and brine of the individual stages, be recovered?

Specific cost [€/a]

Thermal Desalination Processes

65

Sum (thermal energy + capital cost)

Thermal energy

1

Capital cost

Nopt

N

Number of stages [-]

Figure 2.7 Annual specific energy and capital cost vs. number of stages of an MED plant.

It is no coincidence that both questions are linked, because the following paragraph will answer both questions at the same time. The starting point is the schematic of a single-stage evaporation plant with energy recovery (see Figure 2.3). Figure 2.8 shows the logical extension of this single-stage nature into a multiple-effect plant. This arrangement is known in the expert jargon as a “counter-current process”, as the feed flows countercurrently to the distillate and brine. The characteristic feature of this arrangement is one preheater and one evaporator in each stage. Additionally one external preheater and one external final condenser are required. “External” because they are not incorporated into the stages and heated or cooled by an external energy source and therefore usually to be designed separately. As in the design for a single-stage plant, the external preheater exists in order to heat up the feed flow to evaporation temperature. Hence the terminal temperature difference and the boiling point elevation must be bridged. According to Eq. 2.14 this means:

QH ,PH ,1

mF cP ( TTTD

TBPE )

(2.20)

mF

mB1

mD1

mB2

2

mD2

Figure 2.8 Schematic representation of a counter-current MED plant.

Condensate

Heating steam

1

mB3

3

mD3

mBW

N

Hot steam

Brine

Seawater as feed Distillate

mB

Distillate

Brine reject

Seawater in

66 Desalination 2nd Edition

Thermal Desalination Processes

67

The heat necessary for the evaporation in the 1st stage is given by Eq. 2.17 as follows:

QH , E

mD ,1

hV ,TV ,1

(2.21)

The vapour, which is produced in Stage 1 condenses at the surface of the 1st stage preheater and also in the 2nd stage condenser/evaporator unit and in turn generates vapour, whose volume, however, is reduced by the amount of the flow, which condenses in the preheater: the multiple-effect principle explained above, by which the distillate flow is equal to the product of the heating steam flow multiplied by the number of stages, is thus not 100% correct. The solution to this dilemma is that the energy stored in the brine and in the distillate can be utilised. In order to demonstrate this we have to follow the path of both of these flows through the plant. In a multiple-effect plant, the brine and the distillate are fed from stage to stage. Both brine and distillate are leaving the stage as “boiling water”. The boiling temperature and the boiling pressure are determined by the conditions in each stage. To transfer the heat in the condenser/ evaporator unit a driving temperature difference must be present. This means that both brine and distillate enter a chamber in which a lower pressure and a lower boiling point temperature exist. The only possibility to release this surplus energy present under these conditions is through spontaneous vapour production. This process is known as “flash evaporation” or “flashing”. The process is illustrated systematically in Figure 2.9 in an enthalpyentropy (h-s) diagram. The flashing process is an isenthalpic (h=constant) between the pressure levels and . Since state lies in the two-phase region (liquid/vapour) saturated steam and boiling liquid are produced. Figure 2.10 shows two stages of an evaporation plant with the relevant temperatures and mass flows. If boiling point elevation is neglected the mass and energy balance leads to two equations for the mass flows produced by brine and distillate flashing:

mB ,Flash,i

hV ,TV ,i

mB ,i 1 cP , B (TV ,i 1 TV ,i )

mB ,Flash,i

cP , B (TV ,i 1 TV ,i )

mB ,i 1

hV ,TV ,i

mD ,Flash,i

hV ,TV ,i

cP , B

Tstage ,i

(2.22)

hV ,TV ,i

mD ,i 1 cP ,D (TV ,i 1 TV ,i )

mD ,Flash,i

cP ,D (TV ,i 1 TV ,i )

mD ,i 1

hV ,TV ,i

cP ,D

Tstage ,i hV ,TV ,i

(2.23)

68

Desalination 2nd Edition

CP Critical point

Enthalpy h [kJ/kg]

3

Saturated steam

p1 CP p2 1

2

Two-phase region

4 Boiling fluid

Entropy s [kJ/kgK]

Figure 2.9 Enthalpy-entropy (h-s) diagram of a flash evaporation process.

ΔTStage

i–1

i

TV,i–1

TV,i mD,Flash,i mB,Flash,i

mB,i–2

mB,i–1

mB,i

mD,i–2

mD,i–1

mD,i

Figure 2.10 Schematic representation of the flashing process.

The difference in the boiling temperatures between the stages is known as the stage decrement ΔTstage and can be determined from the overall temperature difference of the whole plant ΔT0 and the number of stages as follows:

Tstage

T0 N

Thermal Desalination Processes

69

The temperature profile of an MED plant is shown in Figure 2.11 with all relevant temperature differences. The vapour produced by flashing is added to the vapour produced by evaporation, so that the total distillate produced in a stage comprises 3 parts:

mD ,i

mD , E ,i mD ,Flash,i mB ,Flash,i

(2.24)

Assuming that the stage decrement is for instance 3.5 K and taking into consideration a stage with an evaporation temperature of 40 °C, the ratio of vapour produced by flashing to the brine flow into stage i may be calculated from Eq. 2.22 as follows:

mB ,Flash,i mB ,i 1

cP , B

Tstage ,i hV ,TV ,i

4. 0

kJ 3.5K kg K kJ 2, 407 kg

0.0058

0. 6 %

This means that only about 0.6% of the brine evaporates during the flash process. This amount is relatively small with regard to the total mass balance, that is to say the mass of distillate produced, however the vapour produced by flashing is important for the energy balance. While the vapour generated in the evaporation unit mD , E ,i condenses in the condenser/evaporator unit of the succeeding stage, the flash steam mD , Flash,i mB , Flash ,i condenses on the surface of the stage preheater. So that steam is also available for the preheating of the feed in the 1st stage, the feed flow must be preheated by the stage decrement to be able to flash at the inlet into the 1st stage. This can take place in the external preheater. The share of the heat needed for this process is given by:

QH ,PH ,2

mF cP

TStage

(2.25)

If all the heat flows which are necessary for the process are added together this gives:

QH

QH , E QH ,PH ,1 QH ,PH ,2

(2.26)

According to the multiple-effect principle the distillate mass flow generated by evaporation in the individual stages may be calculated approximately from the amount of total distillate and the number of stages. As the volume of distillate produced by flashing is relatively small (see the sample calculation above), this is permissible.

ΔTT TD

ΔTStage

out

Preheater

THS ΔTHT TTTD ΔTBPE TR ΔTStage T0 TSW

in out

1

in out 2

in out

temperature of the heating steam temp. difference for heat transfer terminal temperature difference boiling point elevation temperature of rejected seawater temperature difference between two stages overall temperature difference temperature of the seawater

ΔTHT

3

ΔTHT

Stages

in out

Figure 2.11 Ideal temperature profile of an MED plant (counter-current flow).

ΔTBPE

THS

Temperature

4

ΔTHT

in out 5

ΔTHT

in out 6

ΔTHT

in out 7

ΔT0

in

TR TSW

70 Desalination 2nd Edition

Thermal Desalination Processes

71

Thus it is possible to arrive to an approximation for the specific heat demand of a multiple effect distillation plant with counter-current flow arrangement:

QH mD

hV ,TV ,m

CF c ( TStage CF 1 P

N

TTTD

TBPE )

(2.27)

The share which is due to the preheating may be calculated in an example as follows: Stage decrement: ΔTstage = 3.5 K Terminal temperature difference: ΔTTTD = 2.0 K Boiling point elevation: ΔTBPE = 0.8 K Concentration factor: CF = 1.4

QH ,PH mD

kJ 1.4 4.0 (3.5 2.0 0.8)K kgK 1.4 1

88.2

kJ kg

The share which may be ascribed to the evaporation was previously calculated above as:

QH , E mD

hV ,TV ,m

2, 407

N

kJ kg

240.7

10

kJ kg

Thus the total heat demand is:

QH mD

240.7

kJ kJ 88.2 kg kg

328, 9

kJ kg

91.36

kWh to

The proportions for the evaporation and for the preheating in relation to the total energy demand are 73% and 27% respectively. Occasionally in the literature a rule of thumb is used, which does not explicitly formulate the preheating but rather takes it into consideration as an exponent for the number of stages, i.e.:

QH , E

hV ,TV ,m

mD

N 0.85

(2.28)

If we insert the values given in the example above, we arrive at

QH , E mD

2, 407 kJ 100.85 kg

334

kJ kg

92.78

kWh to

72

Desalination 2nd Edition

The error is smaller than 5%, so that the rule of thumb is sufficient for a quick calculation, particularly as the energy demand is estimated conservatively, i.e. larger than in the exact calculation. However, on no account should any calculation be made using a formula which uses the number of stages without an exponent, because, as may easily be seen, the result would be 240.7 kJ/kg and thus more than 26% too low compared to the exact calculation, which includes also the preheating. At this point, before the mass and energy balance for an MSF plant is discussed in the following, the determination of the required heat exchanger area should be shown. We thus return to the optimisation task formulated above regarding the optimum number of stages for an MED plant. It is necessary to calculate the capital costs first. These in turn are directly proportional to the costs for the heat exchanger areas, so that if the number of square metres required is known, this can be projected onto the capital costs [1]. The total heat exchanger area of an MED plant can be divided into 3 parts: the preheaters, the evaporators and the final condenser:

AME

APH

AE

AC

(2.29)

The calculation procedure may start with the following well-known general equation for the heat transfer:

Q

k A

THT

(2.30)

Whereby the overall heat transfer coefficient k and the driving temperature difference ΔTHT must be selected properly for each individual case. The steam coming from the previous stage condenses in the evaporator of the succeeding stage and produces the same amount of steam, i.e. the heat transferred in the evaporator unit of stage i may be expressed as:

QE ,i

mD ,i

hV ,i

(2.31)

If Eq. 2.30 is used to eliminate the heat flow Q in Eq. 2.31, then the following term is given for the evaporator area of a stage i:

AE ,i

mD ,i

hV ,i

kE ,i

THT ,i

(2.32)

Thermal Desalination Processes

73

If it is assumed that the evaporator areas of the individual stages are equal in size, the total area can be calculated by multiplication of the above equation by the number of stages N:

AE

N AE ,i

N

mD ,i

hV ,i

kE ,i

THT ,i

(2.33)

Using the following steps this equation may be further simplified: 1. A mean value over the number of stages for the specific evaporation heat and the heat transfer coefficient is introduced:

hV ,i

hV ;

kE ,i

kE

2. The product of the number of stages and the amount of vapour per stage is equal to the total distillate volume (i.e. disregarding flashing):

N m D ,i

mD

3. The driving temperature difference for the heat transfer is derived from the stage temperature difference minus boiling point elevation and is the same for every stage:

THT

THT ,i

TStage ,i

TBPE ,i

T0 N

TBPE

Thus, a simple formula can be derived for the specific evaporator area:

AE mD

kE

hV THT

N hV kE ( T0 N TBPE )

(2.34)

The temperature profile in Figure 2.11 and a temperature vs. heat transfer area diagram such as Figure 2.12 can be used for the determination of the area of a preheater. The feed flow is heated up by the stage temperature difference in one preheater:

QPH ,i

mF cP ,i

Tstage ,i

(2.35)

74

Desalination 2nd Edition T Vapour

TV ΔTTTD

Mak

e-up

ΔTStage

ΔTHT + ΔTBPE

feed

100 %

0 Preheater area

Figure 2.12 Temperature vs. heat transfer area for a preheater in an MED plant (counter -current flow).

With the help of Eq. 2.30 the heat transferred may be explicitly formulated in relation to the area. In this case the mean logarithmic temperature difference must be used as the driving temperature gradient.

QPH ,i

kPH ,i APH ,i

THT ,i

kPH ,i APH ,i

Tln,i

(2.36)

The mean logarithmic temperature difference for a stage i is determined according to Figure 2.11 as:

Tln,i ln

TTTD ,i

TStage ,i TBPE ,i TTTD ,i

THT ,i

(2.37)

The area of a preheater can now be determined from Eqs. 2.35 and 2.36:

APH ,i

mF cP ,i kPH ,i

Tstage ,i

(2.38)

Tln,i

If this equation is multiplied by the number of stages, then the total preheater area of an MED plant (counter-current flow) is derived as:

APH

N APH ,i

N mF cP ,i kPH ,i

Tstage ,i Tln,i

(2.39)

Thermal Desalination Processes

75

The equation can now be simplified similarly to the procedure for the evaporator area. In addition the relationship between the feed flow, the distillate flow and the concentration factor as per Eq. 2.5 can be used.

APH

mD

c CF N P kPH CF 1

TStage ,i

(2.40)

Tln,i

If the mean logarithmic temperature difference as per Eq. 2.37 is now introduced, the specific preheater area is computed as:

c CF N P ln 1 kPH CF 1

APH mD

TBPE THT TTTD

(2.41)

or

APH mD

c CF N P ln 1 kPH N CF 1

T0 TTTD

(2.42)

It should be noted that the external preheater, which is supplied by heating steam from the boiler, is also included in this equation through the definition of the overall temperature difference ΔT0, but not the final condenser. When designing the final condenser the varying temperature of the seawater in the course of the year must be taken into account. Normally allowance is made for the change in the seawater temperature through a change in the mass flow, which flows through the final condenser as a coolant. Figure 2.13 shows a typical flow chart. The layout can be made on the basis of the following calculation scheme:

QC

mSW cP,SW

AC

TSW,out TSW,in

mSW cP,SW kC

TSW,out TSW,in Tln

with

Tln

TSW,out TSW,in ln

kC AC

Tv TSW,in Tv TSW,out

Tln

76

Desalination 2nd Edition Make-up feed mF Vapour mV,N

mRej

Reject

Final condenser Vapour/ liquid separator

mSW

Seawater

mD,N Distillate

Figure 2.13 Simplified flow sheet of the final condenser.

Thus the equations for the determination of the heat exchanger areas needed in an MED plant (counter-current mode) are given. For an estimation of the capital costs in the preliminary design phase it is of assistance if the set of formulae is further simplified. This takes place in a way similar to that for the specific energy consumption, through taking into account the preheating by a correction factor. If the areas are calculated, which are required for the evaporation, then this results in more than 90% of the total area being required for the evaporation and the preheating (including the final condenser). The formula can therefore be simplified as:

AME

AE

AME

APH

AC

1.1 AE

Thus, the following equations may be used for a quick and simple determination of the specific energy consumption and the specific area of an MED plant (cf. Eqs. 2.28 and 2.34):

AE mD

1. 1

kE

QH , E

hV ,TV ,m

mD

N 0.85

hV THT

1. 1

N hV kE ( T0 N TBPE )

(2.28a)

(2.34a)

2,500

1,250

2,000

1,000 Area

1,500

750 Heat

1,000

500

500

250

0

0

5

10 Number of stages N [-]

15

20

77

Specific heat transfer area AE/mD [m2/kg/s]

Specific heat consumption Q/mD [kJ/kg]

Thermal Desalination Processes

0

Figure 2.14 Specific heat consumption and specific area of an MED plant as a function of number of stages.

Figure 2.14 shows the specific heat consumption and the specific area as a function of the number of stages for the following sample values: Overall temperature difference: Heat of evaporation:

ΔT0 = 30 K hV

2, 300

kJ kg

kW m2 K

Overall heat transfer coefficient:

kE

Boiling Point Elevation:

ΔTBPE = 0.7 K

3. 0

The different shape of the two curves (with an increased number of stages the heat consumption decreases and the area increases) indicates that there must be an optimal number of stages when the sum of the energy costs (directly related to the heat consumption) and the capital costs (directly related to the area required) is calculated [1].

mB1

mD1

mB2

2

mD2

mB3

3

Figure 2.15 Schematic of a multiple-effect distillation (MED) as parallel feed flow.

Condensate

Heating steam

1

mD3

N

mDN

Seawater in

Steam

Distillate

Brine

Distillate, mF

Brine, mS

Seawater in, mF

78 Desalination 2nd Edition

Seawater

Brine

. mB

Figure 2.16 Flow sheet of a Multi-Stage-Flash (MSF) evaporation plant. (Once-through mode).

Heating steam . QH

Drinking water storage

. mD

Distillate . mF

Seawater

Feed

Thermal Desalination Processes 79

80

Desalination 2nd Edition

Another type of an MED plant is called “Parallel feed flow” which is shown in Figure 2.15. Here the characteristic feature is that each stage comprises a horizontal heat transfer tube bundle. The heating steam is added to the 1st effect from the boiler and the condensate from the 1st effect is sent to boiler again as shown. This is termed as “Parallel feed flow” because the feed water is distributed equally in parallel to all effects.

2.2.3 Multi-Stage-Flash Evaporation The flow sheet of a Multi-Stage-Flash (MSF) evaporation plant in the socalled once-through mode is illustrated in Figure 2.16. In this mode the feed flows through the preheaters and enters the first stage where the flash process starts. The flashing progresses from stage to stage. The vapour generated condenses on each of the preheaters; the distillate is collected, led into the next stage and removed from the last stage as a product. The concentrated brine flow is rejected from the last stage into the sea. An MSF plant may also be operated in the so-called brine-recycle mode. This will be discussed in more detail later. The kind of operation mode is not significant for the calculation of the heat, which has to be supplied. From the flow sheet it may be recognised that only one external preheater is present in an MSF plant. This preheater is usually designated as a brine heater. An externally heated evaporator or evaporators in the individual stages are not present – in contrast to the multiple-effect plant. Figure 2.17 shows the temperature profile of an MSF plant across the stages. The seawater enters the final condenser, absorbs the heat of the condensing vapour from the last stage. A part of the incoming seawater is rejected back into the sea. The remainder, the so-called make-up feed or simply make-up flows through the preheaters of the stages to the top of the plant. In principle this is a long tube in which the feed is successively heated up through the condensing vapour produced by brine flashing in the chambers. In the brine heater the feed flow is then raised to the temperature level required. An energy balance for the brine heater gives the following result for the heat flow, which needs to be supplied: QH

mF cP ( TStage

TTTD

TBPE

TLosses )

(2.43)

As with the MED plant several temperature differences must be considered in MSF: ΔTStage Stage temperature difference or stage decrement, in order to have a driving force for the flash evaporation in the 1st stage.

Figure 2.17 Ideal temperature profile of a Multi-Stage-Flash evaporation plant. (Once-through mode).

Thermal Desalination Processes 81

82

Desalination 2nd Edition

ΔTTTD Terminal Temperature Difference, because a temperature gradient must exist at the preheater outlet in order for the heat to be transferred ΔTBPE Boiling Point Elevation, because of the salt content of the seawater ΔTLosses Losses, because as a result of the finite length of the chamber the flash process cannot continue until equilibrium is reached (socalled non-equilibrium losses) The temperature is reduced from stage to stage by flashing, by boiling point elevation and by non-equilibrium losses. The brine leaves the plant at the temperature of the last stage. The difference in the highest temperature at the inlet of the first stage, the so-called Top Brine Temperature and the brine temperature of the final stage is designated as the Overall Temperature Difference: TO

TTOP TB , N

(2.44)

If we divide the Overall Temperature Difference by the number of stages we obtain the stage temperature difference or stage decrement:

TStage

TO N

(2.45)

As derived in Eq. 2.22 the mass of vapour produced by flashing, that means the distillate produced, can be calculated from the mass flow of the boiling liquid, the temperature difference from stage to stage and the heat of evaporation. For the whole plant we obtain:

mD mF

cP

TO

(2.46)

hV ,Tm

If this equation is inserted into Eq. 2.43, the specific heat demand becomes:

QH mD

hV ,Tm cP

TO

and also with Eq. 2.45:

cP

TStage

TTTD

TBPE

TLosses

Thermal Desalination Processes

hV ,Tm

QH , MSF mD

1 N

N

TTTD

TBPE TO

TLosses

83

(2.47)

For comparison here once more the equation for the specific heat demand of an MED plant as per Eq. 2.27:

hV ,TV ,m

QH , ME mD

N

CF c CF 1 P

TO 1 N N

TTTD

TBPE

(2.48)

TO

As the flashing in an MED plant is only a “by-product” and the size of the chamber is determined by the size of the evaporator, the losses through non-equilibrium are not important. Therefore, such a term as “ΔTLosses” does not occur in MED. Furthermore the stage temperature difference was replaced by the Overall Temperature Difference divided by the number of stages (see also Figure 2.17). The equations for MED and MSF are similar only at first sight. An example can help to be able to recognise the difference. Figure 2.18 shows a comparison for the following given data: ΔTO = (70–35) K = 35 K

Overall temperature difference:

Spec. heat consumption QH/mD [kJ/kg]

10,000

1,000

MSF

MED

100 0

5

10

15

Number of stages N [–]

Figure 2.18 Comparison of specific heat consumption of MSF vs. MED.

20

84

Desalination 2nd Edition

Terminal temperature difference: Boiling Point Elevation: Non-equilibrium losses: Concentration factor: Specific heat capacity: Heat of evaporation:

ΔTTTD = 2.0 K ΔTBPE = 0.8 K ΔTLosses = 0.5 K CF = 1.4 kJ c P 4. 0 kgK hV ,Tm

2, 376

kJ kg

We can see that with the number of stages the MED plant becomes more and more efficient concerning the specific heat consumption in comparison with the MSF plant. In any case this does not mean that it is always the best solution to build an MED plant with many stages, because for finding the best solution a calculation of the economic efficiency, i.e. energy cost + capital cost, has to be carried out (see Figure 2.7). For example, we have to take into account that an effect of an MED plant is much more expensive than an MSF stage due to the evaporator in each stage. Figure 2.19 shows the flow sheet of an MSF plant in brine-recycle mode. This option differentiates itself from the once-through mode as a brine flow is circulated. One of the main advantages of brine-recycle mode is that the costs for the pre-treatment can be reduced. Pre-treatment typically involves degassing and conditioning with anti-scaling and anti-foaming chemicals. While in the once-through process the whole of the feed flow must be pre-treated, in the brine-recycle mode only the so-called make-up flow is treated. Significant cost savings can be achieved. A further advantage is the flexible operation of the plant through the additional degree of freedom given in the recycling. There is no meeting of seawater desalination experts at which the advantages and disadvantages of both of the processes were not controversially discussed. Factually the brine-recycle mode has the following disadvantages: a higher energy demand for the large brine-recycle pump a high number of pumps and pipes with the disadvantage to be susceptible to faults as well as the requirement for maintenance and repair a complex layout with all the associated difficulties The advantages and disadvantages of the once-through process are: high costs for the pre-treatment a low level of flexibility in operation

M Make-uo

DG degasifier

ms. out

Reject

mR

mM DG

Figure 2.19 Schematic of a Multi-Stage-Flash evaporation plant as brine-recycle mode.

Heating steam QH

Brine recycle flow mR

Brine

mB, R

mB, S

Seawater in/out, brine reject

Heating steam

Brine recycle

Distillate

Distillate mD

Seawater mS, in

Thermal Desalination Processes 85

86

Desalination 2nd Edition reduced thermal efficiency by rejection of hot brine (40 °C) a lower number of pumps and pipes and thus lower maintenance and repair costs a simple and reliable operation overall a higher level of availability

A conclusive evaluation cannot be given at this point. However, when selecting the optimum process in a specific project both alternatives should always be examined along with the prevailing conditions under the heading “MSF”. The specific heat demand of a brine-recycle plant does not differentiate itself from that of a once-through plant. However, the following should be observed: the mass flow passing through the brine heater is the brine recycle flow, based on the temperature profile for the brine-recycle plant shown in Figure 2.20 the recovery and the reject part of the plant have to be used as an effective driving overall temperature difference for the plant. Thus, the following is arrived at with the help of Eqs. 2.45, 2.46 and 2.47:

Brine-recycle mode:

mD

cP

mR

TO

TStage QH , MSF mD

hV ,Tm N

1 N

(2.49)

hV ,Tm

TTTD

TO N

(2.50)

TBPE TO

TLosses

(2.51)

As briefly mentioned above, the brine-recycle plant has an additional degree of freedom through the recycling in comparison with the oncethrough plant. Through this the balances for mass, salt and energy are more complicated. Based on the relationships and representations in the flow sheet (Figure 2.19) the following set of equations is derived: Mass balance Overall mass balance

mS ,in

mS ,out mB,S mD

0

(1)

ΔTTTD

ΔTStage

out

Brine heater

THS TTBT ΔTTTD ΔTBPE TBN TDN ΔT0 TSW ΔTlosses

TTBT

in out 1

in out 2

temperatureoftheheatingsteam topbrinetemperature terminaltemperaturedifference boilingpointelevation temperatureofthebrineleavingstageN temperatureofthedistillateleavingstageN overalltemperaturedifference temperatureoftheseawater non-equilibriumlosses in out

ΔTStage

3 Stages

in out NRec

in out 5

Figure 2.20 Ideal temperature profile of a Multi-Stage-Flash evaporation plant. (Brine-recycle mode).

ΔTlosses

ΔTBPE+

Temperature

THS

in out N

in out

TSW

TDN

TBN

ΔT0=N*ΔTStage

Thermal Desalination Processes 87

88

Desalination 2nd Edition

Mixing Point (1)

mS ,in

mS ,out

mM

Mixing Point (2)

mB, R mM

Mixing Point (3)

mB, N

(2) in (1) (3) and (6) →

m D m M m B ,S m D m R m B, N

mR

mB,S

0

(2)

0

mB, R

(3) 0

(4) (5) (6)

Material balance (salt balance) from (5) →

0 mM x M

from (6) →

0 mR x R

mB,S x B, N mB, N x B, N

Two different concentration factors may be defined from these equations:

x B ,N

CFM

xM x B ,N

CFR from (7) →

xR mM

mM

(7)

mR mB ,N

(8)

CFM mB ,S

CFM (mM

mM

from (6) →

mM mB ,S

mD )

CFM m CFM 1 D

(9)

The relationship of the make-up mass flow to the distillate flow is designated as the “make-up factor”, i.e.:

mM mD

CFM CFM 1

Energy balance Brine heater

QH

mR cP

TBH

Flashing

mD

hV ,Tm

mR cP

(10)

(11)

T0

(12)

Thermal Desalination Processes

89

Thus a set of equations is now available, with the help of which a complete balance for a MSF plant in brine-recycle mode is possible. The procedure is as follows: Given: Distillate mass flow = product m ˙D Salt concentration of the make-up flow = salt conxM = xSW centration of the seawater Maximum concentration in the final stage xB,N (Given by the solubility of the salts) Overall Temperature Difference ΔTO Top Brine Temperature ΔTTBT Solution algorithm: mD

hV ,Tm

1.

Eq. 12

2.

Eq. 7

3.

Eq. 9

4.

Eq. 10

5.

Eq. 5

mB,S

mM

mD

6.

Eq. 6

mB, N

mR

mD

7.

Eq. 4

mB, R

8.

Eq. 8

xR

mR

cP

x B,N

CFM

mM

TO

xM

CFM m CFM 1 D

mM mD

mB, N

x B,N

mB,S

mB , N mR

In the following section the required condenser area for an MSF plant will now be calculated. Because, analogous to the MED plant, an optimal

90

Desalination 2nd Edition

number of stages must also be derived for an MSF plant. However, this can only be done by analysing the costs and for this the required area must be known. The equations for the heat transfer in a preheater/condenser in an MSF plant are the same for those in an MED plant. This results in the following procedure in line with Figure 2.12:

QMSF ,i QMSF ,i

mD ,i

hV ,i

kMSF ,i AMSF ,i

(2.52)

Tln,i

(2.53)

From this the condenser area for a stage i is given as:

AMSF ,i

mD ,i

hV ,i

kMSF ,i

Tln,i

(2.54)

The condenser area of the whole plant (without the brine heater) can be calculated by multiplying Eq. 2.54 by the total number of stages:

AMSF

N AMSF ,i

N mD ,i kMSF ,i

hV ,i Tln,i

(2.55)

This equation may be further modified in the following steps: 1. A mean value for the heat of evaporation and the heat transfer coefficient over the stages is introduced:

hV ,i

hV ;

kMSF ,i

kMSF

2. The product of the number of stages and the amount of distillate per stage is equal to the total mass of distillate:

N mD ,i

mD

3. The mean logarithm temperature difference to be used for the heat transfer is the same for every stage i (cf. Eq. 2.37):

T0

Tln N ln 1

N

T0 TTTD

(2.56)

Thermal Desalination Processes

91

With these assumptions, the total specific condenser area of an MSF plant is:

AMSF mD

ln 1

hV kMSF

N

T0 TTTD

N

(2.57)

T0

The area required in an MSF plant can now be put in relation to the energy demand. For this reference is made to Eq. 2.51. This reads, slightly modified:

QH , MSF mD

hV ,Tm

TTTD

1 N

TBPE T0

TLosses

(2.58)

The so-called Gained Output Ratio (GOR) is defined as a sort of efficiency measure of a thermal seawater desalination plant. It is assumed that the energy supply for the plant uses steam from a boiler or a turbine. The GOR is defined as the ratio of the distillate produced to heating steam provided:

mD mHS

GOR

If an average heat of evaporation is assumed then the heat introduced into the MSF plant may be expressed as:

QH , MSF

mHS

hV ,Tm

(2.59)

If this equation is entered into Eq. 2.58, the following term for the GOR is computed:

GOR

1 N

TTTD

TBPE TO

TLosses

1

(2.60)

As well as the number of stages N the equations for the specific area (Eq. 2.57) and for the GOR (Eq. 2.60) contain the terminal temperature difference ΔTTTD as a free parameter: the system consists therefore of two equations with two unknowns and can be solved using a substitution procedure, resulting in GOR as a function of the specific area and the number of stages:

GOR

f N,

AMSF mD

Desalination 2nd Edition

92

14 N = 50 Gained output ratio GOR [-]

12

N = 30 N = 20

10

N = 15 N = 10

8

N=5 6 4 2 0 0

50 100 Specific heat transfer area [m²/kg/s]

150

200

Figure 2.21 Gained Output Ratio of an MSF plant vs. specific condenser area (parameter: number of stages N).

The graph of this function is shown in Figure 2.21 for the following sample values. Overall temperature difference: ΔTO = 70 K kJ Heat of evaporation: hV 2, 340 kg kW Overall heat transfer coefficient: kMSF 3.0 2 m K Boiling Point Elevation: ΔTBPE = 0.48 K Non-equilibrium losses: ΔTLosses = 0.2 K The resultant set of curves has far-reaching consequences for the design of an MSF plant. When a GOR is specified, then a correct solution can be obtained from the diagram at various numbers of stages and various condenser areas. We only can obtain the actual number of stages and the condenser area by a cost calculation.

2.2.4 Multiple-Effect Distillation with Thermal Vapour Compression (MED-TVC) The various possibilities to provide the necessary heat for the evaporation process with each of their advantages and disadvantages will be examined more closely in section 3. If steam at a relatively high temperature and

mDIS

mSS

mSS

Q1

TV,1

Suction steam

mB,1

1

mD,1

Q2

TV,2

mB,2

2

mD,2

Q3

TV,3

mB,3

3

mD,3

Figure 2.22 Flow sheet of an MED evaporation plant with thermal vapour compression (Gross balance).

mMS Back to boiler

Discharge steam

Motive steam

mMS

From boiler

QN

TV,N

N

Distillate

mD,N

Brine

mB,N

mRej mSw

Feed

mF

Thermal Desalination Processes 93

94

Desalination 2nd Edition

pressure level is available then thermal vapour compression (TVC) can be utilised to improve the efficiency of the seawater desalination plant. Figure 2.22 shows the flowsheet of a Multiple-Effect-Distillation plant with thermal vapour compression. As can be recognized from the flowsheet the compressor sucks in a part of the vapour from the last stage, compresses it with the help of the motive steam and adds the mixture to the 1st stage as heating steam. Normally the motive steam will be generated in a steam boiler, which means that the mass flow of motive steam is passed back to the boiler as condensate. The following balances may be derived from the flow sheet: mD (N 1) mD ,i mD , N mSS mDis hV ,TDis mD ,1 hV ,T1

Mass of distillate (total) Energy balance (1st stage)

hV ,TDis

hV ,T1

mD ,1 mDis Mass of distillate 1 stage) mD ,1 mDis mSS mMS Mass of distillate (stage N) mD , N mDis mSS mMS From these equations follows that: st

mD mD

(2.61) (2.62)

(2.63) (2.64) (2.65)

(N 1) mDis mMS mSS

(N 1) (mMS mSS ) mMS mSS mD

N (mMS mSS )

(2.66)

The ratio of motive steam to suction steam is usually called entrainment factor w. The entrainment factor is a characteristic feature of the compressor itself depending upon the various pressure levels and the internal efficiencies of the device. The entrainment factor can be obtained from a performance curve of the compressor usually published by the manufacturers (as an example see Figure 2.28) or calculated by thermodynamics as being done later in this chapter. Assuming that we do know the entrainment factor, Eq. 6.66 may be transformed as follows: w

Entrainment factor Total mass of distillate

mD

mSS mMS

N mMS (1 w )

(2.67)

Thermal Desalination Processes Specific mass of motive steam

mMS mD

1 N (1 w )

95

(2.68)

If the entrainment factor and the number of stages of the MED plant is known the amount of motive steam necessary for the compression can be calculated from this equation. As the mass of motive steam is provided by the steam boiler this figure may be equated to the thermal energy demand of the MED – TVC process. Thermal vapour compression is an open process: the motive steam, which comes from the boiler mixes with the suction steam; the mixture is fed into the 1st stage evaporator and appears after the condensation as liquid distillate. A part of this distillate returns to the boiler as boiler feed water, while the rest is added to the distillate of the remaining stages, in other words it becomes “product”. If this product is utilised as drinking water there is the risk that the chemicals, which were added to the boiler feed water and which in this way can end up in the product, pollute the drinking water and make it unfit for human use. Should this be the case, and if no other chemicals are able to be used, which are harmless to man, then the part of the condensate from the 1st stage, which is not returned to the boiler must be rejected or otherwise used. For this special case the following balances may be derived from the flowsheet in Figure 2.23: Total mass of distillate: mD

(N 1) mD ,1 mD ,N

mD

(N 1) mDis mMS

mD

(N 1) (mMS mSS ) mDis

mD

mMS (N 1) 1

mD

m MS (N 1) (1 w ) 1)

mD

m MS N (1 w )

mSS mMS

1

w N

Specific mass of motive steam:

mMS mD

1 N (1 w )

w N

(2.69)

Dump to sea

mSS

mDIS

mSS

Q1

TV,1

mB,1

1

mD,1

Q2

TV,2

mB,2

2

mD,2

Q3

TV,3

mB,3

3

mD,3

Figure 2.23 Flow sheet of an MED evaporation plant with thermal vapour compression (Net balance).

mMS

Discharge steam

Motive steam

mMS

From boiler

QN

TV,N

N

mD,N Distillate

mB,N Brine

mRej mSw

Feed

mF

96 Desalination 2nd Edition

Thermal Desalination Processes

97

w N appears here, which increases the motive steam but whose influence becomes smaller as the number of stages is increased. In the above discussed case vapour is sucked up from the final stage N. This choice is already the result of an optimisation process taking into consideration the thermodynamical efficiency of the compressor chosen. Assuming that the compressor sucks up vapour from any stage NTVC the balances presented above will change as follows: In comparison with the “gross balance” (Eq. 2.69) the term

Gross balance Total mass of distillate

mD mD mMS mMS

Specific mass of motive steam:

mD

Net balance Total mass of distillate

mD mD mMS

Specific mass of motive steam

m MS mD

N NTVC w 1 N NTVC w N m MS N

(2.70)

NTVC 1 mSS

NTVC

1 w

1 N (NTVC

Supersonic nozzle

Mixing chamber

N mM S NTVC mSS

(2.71)

1) w

Diffuser (venturi tail)

Motive steam mMS

1

mMS vMS,out 2

pMS vMS,in Suction steam (load)

0

mSS pSS v=0

Figure 2.24 Scheme of a thermal vapour compressor.

vDiff,in 3

mDIS pDIS vDiff,out 4 Discharge steam

98

Desalination 2nd Edition

In order to determine the mass flow of motive steam as a measure for the energy consumption of an MED/TVC - plant the entrainment factor w must be known. For this it is necessary to discuss thermal vapour compression more in detail. Figure 2.24 shows the schematic construction of a thermal vapour compressor. In order to explain the thermo- and hydro dynamical processes three zones have to be considered: 1. Nozzle 2. Mixing chamber 3. Diffuser We can understand the basic principle of the compressor using the well-known Bernoulli equation, which states that for incompressible fluids the sum of pressure energy, kinetic energy and potential energy along a streamline remains constant:

p

2

v2

g z

const.

(2.72)

At first we should discuss the diffuser shown schematically in Figure 2.24. Assuming that the potential energy between state and remains constant (horizontal arrangement), Eq. 2.72 may be transformed to:

p1

2

v12

p2

2

v22

(2.73)

2 1

A1

A2

min

mout

v1 p1

p2

Figure 2.25 Schematic representation of a diffuser.

v2

Thermal Desalination Processes

99

With the help of the equation of continuity for the mass flow entering and leaving the diffuser, the pressure at the diffuser outlet may be derived as follows:

min

mout

1

v1 A1

2

v2 A2

(2.74)

With ρ = ρ1 = ρ2 we obtain

p2

p1

2

v12

1

A1 A2

2

(2.75)

This equation means in words: If the cross-section A2 is larger than A1, then the pressure p2 is greater than p1! If the cross-section A2 is smaller than A1, then the pressure p2 is lower than p1! or If the velocity decreases in the diffuser, then the pressure increases! If the velocity increases in the diffuser, then the pressure decreases! With the help of these statements the processes in the compressor may be described as follows: in the nozzle the motive steam, which has a high pressure and a relatively low velocity accelerates; at the nozzle outlet the motive steam has a very high velocity, whereby the prevailing pressure in the immediately surrounding region is very low; as a result of this low pressure vapour from the evaporation plant is sucked up into this part of the compressor (usually this vapour is called “suction steam” or “load”); in the mixing chamber motive steam and suction steam mix together; in the diffuser the velocity of the steam mixture is gradually lowered, whereby the pressure increases again. Figure 2.26 shows a characteristic velocity and pressure profile in a thermal compressor [2]. Two phenomena are of particular importance: 1. The velocity at the nozzle outlet is supersonic! 2. In the diffuser the velocity is reduced by a shock wave (“sonic boom”).

100

Desalination 2nd Edition

The Bernoulli equation explains very well the functioning of the TVC process, but the law is only valid for incompressible fluids. Since steam is compressible the use of the Bernoulli equation is not quite correct for an energy balance around the nozzle, the mixing chamber and the diffuser. The 1st law of thermodynamics and the principle of momentum transport must be drawn upon. Furthermore steam behaviour is not described accurately by the ideal gas law for the operation conditions in such ejectors. Calculations are best done by using an enthalpy vs. entropy (h-s) - diagram (see Figure 2.27). The different states are mainly defined by three pressures: Motive pressure (steam produced by the steam boiler)

pMS

Mixing Chamber

Nozzle Motive steam mMS pMS vMS,in

Diffuser Normal shock

3 mMS 1

mDis

2M v MS,out

vDiff,in

2S

4

pDis vDiff,out

0 Discharge

Evaporator connection

mSS pSS

Suction

v=0 2M

Velocity

Motive

3

Sonic 2S 1

4

Load

0

pDis

1 Pressure

Motive 4

3 Load pSS =pBP,NTVC

2

0

Figure 2.26 Pressure and velocity profile in a thermal vapour compressor.

pDis

Thermal Desalination Processes pSS = pBP,NTVC

101

Suction pressure or load pressure (steam pressure in the suction stage of the MED/TVC plant) Discharge pressure (condensation pressure in the 1st stage evaporator)

pDis

If the temperature profile for an MED plant is given as in Figure 2.11, then the various pressures may be derived from boiling pressure data. Since the motive steam pressure and the motive steam temperature is usually given by the performance of the steam boiler, the starting point 1 in the h-s diagram is also defined. In the nozzle (1 2) the velocity increases and the pressure decreases down to suction pressure pss. This change of state is polytropic. The mixing point 3 is located on the line of constant pressure pss. To obtain the exact place of point 3 on this line, point 1 and point 0 can be connected by a straight line as to motive steam and suction steam are mixed up. The mixing point 4* is located at the crossing of the pressure line pDis and the mixing line. 4* leads to 3* but not to 3 or 3´! Due

S

pM

1

sirr h4

4 4 *

4’

Enthalpyh [kJ/kg]

x=1 0 s p Di

3*

3

p SS

MS SS Dis

2 2’

Motive steam Suction steam Discharge steam

Entropy s [kJ/kgK]

Figure 2.27 Enthalpy vs. entropy (h-s) - diagram for the TVC process.

Desalination 2nd Edition

102

to the irreversibilities in the nozzle and the diffuser the changes of state are not isentropic. A solution is given by the introduction of efficiencies for the nozzle and the diffuser. Solving the energy balance for the different parts of the compressor leads to the following equation:

mSS mMS

irr ,noz

irr ,dif

h1

h2

h4

h3

1

(2.76)

Since the isobars are nearly parallel (see h-s – diagram), it can be assumed that h4

h3

h4*

h3*

and this leads to the entrainment factor w:

mSS mMS

w

irr ,noz

irr ,dif

h1 h2 1 h4* h3*

(2.77)

To get familiar with this procedure we would like to calculate the following example: Given: Motive pressure Suction pressure Discharge pressure Efficiency of the nozzle Efficiency of the diffuser

pMS = p1 = 10 bar pSS = p0 = 0.074 bar pDis = p4 = 0.2 bar ηirr,noz = 0.9 ηirr,dif = 0.7

saturated steam TV = 40 °C TV = 60 °C

Solution: (by using the h-s diagram)

h1

2, 776.1

w

kJ h kg 2

0. 9 0. 7

2, 060.0

kJ h kg 4*

2, 610.0

kJ h kg 3*

2, 470.0

kJ kg

2, 776.1 2, 060 1 0.795 2, 610 2, 470

With this figure, the specific steam consumption of a 4 - staged TVC – plant can be calculated as follows (gross balance):

Thermal Desalination Processes 20 15 12 10

Rs = kg motive steam per kg load

8 7 6 5.0

Discharge pressure/suction pressure

103

4.0 3.5 3.0 2.6 2.4 2.2 2.0 1.8 1.7 1.6 1.5 1.4 1.3 1.20 1.16 1.14 1.12 1.10 1.09 1.08 1.07 1.06 1.05 1.6 1.7 1.8 1.9 2 2.2

2.6

3

4

5 6

8

10

15 20 30 50

2

10

2 5 5 103 104 6

Motive pressure/suction pressure

Figure 2.28 Performance diagram for a thermal vapour compressor [2].

mMS mD

1 N (1 w )

1 4 (1 0.795)

1 7.8

0.139

The Gained Output Ratio GOR of such a plant is given by: mD GOR N (1 w ) 7.8 mMS Data sheets may be found in the literature or produced by manufacturers with the help of which it is possible to determine the entrainment factor w at given pressures. Figure 2.28 shows such a diagram for a thermal compressor which is in principle similar to a performance curve for centrifugal pumps (Δh vs. V) [2]. For the above discussed example we obtain:

pDIS pSS pMS pSS

2.7 135

104

Desalination 2nd Edition

From diagram (Figure 2.28) we get

w

1 RS

1 1.25

0. 8

and according to the formula used for calculating the GOR we finally obtain:

GOR

mD mMS

N (1 w ) 7.2

The difference between this value and the more exactly calculated value using the h-s-diagram is less than 10% (7.2 to 7.8) Therefore it is appropriate using a performance diagram for rough calculation of the GOR.

2.2.5 Single-Stage Evaporation with Mechanically Driven Vapour Compression In this desalination process the thermo-compressor described in the previous paragraph is replaced by a mechanically driven vapour compressor. With regard to the energy supply it is a process, which requires solely mechanical energy and thus is comparable with the reverse osmosis. However, the actual desalination process is of a thermal nature (separation of salt and water using evaporation), whereas in reverse osmosis the separation of salt and water takes place as a result of a trans-membrane pressure difference and is therefore purely mechanical. Figure 2.29 shows the principle of a single-stage evaporation plant with mechanical vapour compression (MVC) and a picture of a typical plant. The process is essentially composed of the following elements: Evaporation unit Vapour compressor Feed – distillate – preheater (1) Feed – brine – preheater (2) Circulating pump Externally heated preheater (3) for the circulation flow In order to be able to better understand the thermodynamics of mechanical vapour compression, a representation in the h-s diagram is helpful (see Figure 2.30). Part of the water, which is usually sprayed into

Thermal Desalination Processes

105

mF Compressor

mD 2

M 1

Evaporator

3 mD + minj 4 5

Preheater 3

minj

Preheater 1

mF Feed

Preheater 2 mB

mD

Brine

Distillate

Figure 2.29 Flow sheet and picture of an evaporation plant with mechanical vapour compression (by courtesy of VA TECH WABAG GmbH).

Desalination 2nd Edition

106

p3

p2

3

CP Critical point

T3

Enthalpyh[kJ/kg]

4 2

T2

CP

5 1

Entropy s [kJ/kgK]

Figure 2.30 Mechanical vapour compression process in an h-s diagram.

the evaporation chamber is continuously evaporated (1 2). This vapour, which has a certain temperature according to the pressure in the evaporator, is sucked up by the compressor and brought up to a higher pressure (2 3). In the enthalpy – entropy diagram the state 3 after the compression is in the superheated region. Through the injection of a small amount of distillate saturated steam according to the pressure present at the compressor outlet (3 4) is produced. As this pressure, and thus the saturation temperature is higher than that on the outside of the evaporator tube bundle, this vapour can be used 5) takes usually place on the as heating steam. The condensation (4 inside of a horizontal tube bundle. It follows from an energy balance around the evaporation chamber that the condensation of 1 kg of heating steam on the inside of the tubes produces 1 kg of steam outside the tubes. However, energy is still needed in order to preheat the seawater to the evaporation temperature. Heat exchangers are used for this, through which the seawater flows and which absorbs the sensible heat of the distillate and the concentrate. As a driving temperature difference (i.e. a terminal temperature difference ΔTTTD) is

Thermal Desalination Processes

107

needed for the heat transfer, the preheating by the distillate and the brine is insufficient for bringing the feed up to evaporation temperature. Usually the recycle flow is heated up by electric heating. Additionally, the external heating also covers the insulation losses, which amount to about 1% of the evaporation heat. The circulation of a certain amount of water is most necessary in order to be sure that the heat exchanger surfaces, here the outside of the tubes, are always wetted. If the film would break up, at these points salt would crystallise. The scaling would not just hinder the heat transfer, but in the end the tube bundle would also become totally blocked. The mechanical energy introduced into the process is given by the following equation:

Pt

mD

h3 h2

(2.78)

Assuming a polytropic change of state in the compressor, the following equations for the temperatures and the pressures at the inlet and outlet of the compressor are valid: Polytropic change of state: p·Vn = Const. (n = polytropic exponent) Ideal gas law: p·V = R·T p·V = R·T

T3

T2

p2 p3

n n 1

(2.79)

The temperatures and pressures are coupled with the evaporation process as follows: Temperature at compressor inlet = Evaporation temperature Tv T2 p2 Pressure at compressor inlet = Evaporation pressure pv p3 Pressure at compressor outlet = Eondensation pressure pc The condensation pressure is given by the condensation temperature according to the boiling pressure curve. The condensation temperature may be derived from the evaporation temperature plus a driving temperature difference necessary for the heat transfer:

TC

TV

THT

(2.80)

The technical work for a steady-state change of condition may be derived using the 1st law of thermodynamics. This means: 3

wt 23

w friction,23

vdp 2

(eout 2

eout1 )

(2.81)

108

Desalination 2nd Edition

Besides friction and changes in external energy, the integral represents the most important part of the technical work. The integral for a polytropic change of state may be expressed as follows: 3

n (p v p v ) n 1 3 3 2 2

vdp 2

(2.82)

With the help of Eq. 2.80 we finally arrive to: 3

n R (T3 T2 ) n 1

vdp 2

3

n

vdp

R T2

n 1

2

T3 1 T2

n R T2 n 1 p3 p2

n n 1

1

(2.83)

The polytropic exponent n can be expressed as follows:

n n 1

k k 1

VC

with k Isentropic exponent (k = 1.326 for water, T < 100 °C) Efficiency of the compression ηvc Thus Eq. 2.83 becomes: 3

vdp 2

k k 1

VC

R T2

p3 p2

k k 1 vc

1

(2.84)

If the inner friction in the compressor is taken into account using an efficiency factor ηVC and if we ignore the changes in external energies, we obtain the following term for the power demand of the compressor (cf. Eq. 2.78):

Pt Pt

mD

mD h3

k k 1

VC

R T2

h2 p3 p2

k k 1

vc

1

(2.85)

Hence, in order to calculate the power consumption only the capacity of the plant m˙D needs to be known, as well as the pressures and temperatures

Thermal Desalination Processes

109

at which the evaporation and the condensation take place. According to Eq. 2.80 this means that a temperature difference ΔTHT must be fixed for the heat transfer. This is an optimisation problem as this temperature difference appears in two equations: firstly in the heat transfer equation and secondly in the compressor performance. These can be expressed as: 1.

Q

k A

THT

2.

Pt

mD

n R n 1

mD THT

hV ,TV 1

(2.86) (2.87)

VC

The equations above demonstrate that the heat exchanger surface, and thus the costs for the compressor, become smaller the greater the temperature difference and the greater the heat transfer coefficient. When attempting to minimise the investment costs, large temperature differences and large heat transfer coefficients must be aimed for. However, the temperature difference may not simply be selected to have an arbitrary value for two reasons: 1. The temperature difference for the heat transfer is directly proportional to the compressor power consumption according to the equation listed above. If cheap and simply constructed fans were to be employed, the technical limit for a mechanically driven vapour compressor is a maximum of 10 Degree K. Greater temperature differences can only be achieved using expensive screw-type compressors. However, their use is not worthwhile in comparison with the savings in the heat exchange surface. 2. The evaporation of seawater is constrained by the crystallisation of salts (calcium carbonate, calcium sulphate). If bubbles are formed during the evaporation – this process may be observed in any kitchen when water boils – salt crystals will develop at the edges of the bubbles on the evaporator surface and in time will encrust the whole surface. Boiling with bubbles occurs when the heat flow density (Watt per m2) or the driving temperature difference exceeds a specific value. For water or seawater this is at about 7 K this limit should on no account be exceeded in seawater desalination [3]. For the reasons given there are the following constraints on mechanical exhaust vapour compression in practice: 1. A single-stage process is used as only small temperature differences can be overcome by the compressor (as opposed to

condensate

mHS,E

preheater

Q1 m

TV,1

1

Figure 2.31 Schematic of an MED plant and definition of GOR.

steam

mHS

mHS,PH

Q2

TV,2

GOR =

2

mD mHS

Q3

TV,3

3

QN

TV,N

N

mD

mSW

110 Desalination 2nd Edition

Thermal Desalination Processes

111

thermal compression where greater temperature differences and thus also multiple-effect plants are possible). 2. The costs for a compressor are a constraint on the capacity of the plant, i.e. normally only small plants are equipped with mechanical vapour compression (up to approx. 2,000 m3/d).

2.3

Performance of Thermal Desalination Processes

2.3.1 Definition of Gained Output Ratio If a query is made as to the energy efficiency of a thermal seawater desalination plant in most cases the Gained Output Ratio, GOR for short, will be given. The GOR is defined as the ratio of fresh water produced to the heating steam and is in line with the definition of efficiency, i.e. benefits to effort:

fresh water produced

fresh water produced

energy demand

heating steam

GOR

The GOR is thus the relationship of two mass flows, the distillate flow and the heating steam flow, as in the schematic representation of an MED plant in Figure 2.31, which brings out the following units.

GOR

mD mHS

Unit :

GOR

kg/s kg/s

(2.88)

Thermal plants with the same GOR therefore require the same amount of heating steam for the production of a certain mass of distillate. In saying this it is simply assumed that this is saturated steam, there is no information on the exact conditions of the heating steam, means pressure and temperature. Thus, the energy content of the heating steam is not clearly defined. It cannot therefore be concluded that both plants consume exactly the same amount of fuel or that, for instance, the costs for the supply of thermal energy are the same. The GOR is therefore not quite suitable as a scale for a process comparison. For this the energy supply station and the costs for the fuel must also be taken into consideration [1]. The specific heat demand of thermal seawater desalination plants, related to the number of stages and the thermodynamic data, was already determined in section 2.2. As may be seen from Figure 2.31, this heat is equal to

112

Desalination 2nd Edition

the condensation heat of the heating steam. (Sub-cooling of the condensate is not taken into consideration at this point.) The GOR can therefore be calculated from the specific heat demand of the plant as follows:

Q mD

mHS

Q mD GOR

hV ,HS

(2.89)

mD hV ,HS

(2.90)

GOR hV ,HS

mD Q

(2.91)

The last term is also known in the literature [4] as the Performance Ratio (PR).

PR

hV ,HS

hV ,HS

mD Q

Q

(2.92)

mD

The heat of evaporation is usually directly entered in this equation as numerical value with a guideline for the units of the heat and the distillate mass flow. Thus, the equation mutates to a numerical value equation. We strongly recommend to maintain a distance to such definitions, as they create more confusion than clarity as a measure for the energy demand of a thermal plant. As an example of the procedure described above the GOR for an MED plant will be determined in the following text. A rough formula for the specific heat demand of an MED plant is (see Eq. 2.28):

Q mD

hV ,Tm N 0.85

(2.93)

Where Tm is the average stage temperature in MED plant. Eq. 2.91 it follows:

GOR

hV ,HS hV ,Tm

N 0.85

(2.94)

If it is assumed that the evaporators in the different stages of an MED plant are all of the same construction, in particular the evaporator in the 1st stage, then the temperature of the heating steam is given by the temperature

Thermal Desalination Processes

113

profile of the plant. Thus, for a typical MED plant the ratio of the specific heats of evaporation in Eq. 2.94 can be determined as follows: Example: Number of stages,

N

10

Vapour temperature 1st stage

T V,1

650 C

Vapour temperature last effect

TV,N

350 C

Stage decrement

ΔTStage

Heating steam temperature

TV,HS

68.3 °C

Heat of evaporation (MED plant)

ΔhV,Tm

2,382

Heat of evaporation (Heating steam)

ΔhV,HS

2,338

GOR

2, 338 0.85 10 2, 382 GOR

0.98 100.85

6.93

The 2% - error (0.98 to 1.00), which results from the fact that different values for the heat of evaporation have to be taken into account in the GOR equation, is therefore comparatively small. If these points are taken into consideration the gained output ratio can always be used if it is simply a question of obtaining a quick statement on Grid

Brine heater / avaporator 1st cell

HP 1

Fuel

LP Desalination

Air

Combustion Steam boiler >Top brine temperature

Boiler feed water pump 1

Pressure control valve

Figure 2.32 Flow chart of a single-purpose plant (Steam Boiler + Desalination).

114

Desalination 2nd Edition Tin, pin

mHS, in Tin, pin

enthalpy h [kJ/kg]

Δh

minj

Tout, pout

mHS, out Tout, pout entropy s [kJ/kgK]

Figure 2.33 Schematic of the pressure reduction station with injection for de-superheating and respective h-s diagram.

the efficiency of a thermal seawater desalination plant. The GOR is unsuitable as a parameter for a comparative assessment of the energy cost of different seawater desalination processes.

2.3.2

Single Purpose vs. Dual Purpose Plants

Figure 2.32 shows the typical flow sheet for a single purpose plant. Steam is produced in the steam boiler at a relatively high pressure level (HP, 10 to 20 bar). The steam is normally lightly superheated. With the exception of MED plants with thermal vapour compression MSF and MED plants require heating steam at a low temperature and pressure level (LP): MSF → Saturated steam at 2.0 bar/ approx. 120 °C MED → Saturated steam at 0.3 bar/ approx. 70 °C The high pressure of the saturated steam must therefore be reduced before it enters the heat exchanger of the evaporator plant (brine heater for MSF, evaporator of the 1st stage for MED). As shown in Figure 2.33, this happens using a throttle valve in combination with water injection. The latter is necessary in order to transform the steam, which is superheated after the throttle into saturated steam. Otherwise superheated steam would enter the heat exchanger. However, as superheated steam acts like a gas, the heat transfer coefficient on the gas side would be very low and the area of the heat exchanger would have to be increased enormously. This makes neither technical nor economic sense. The injection is a simple and effective

Thermal Desalination Processes

115

method of ensuring that saturated steam flows into the heat exchanger, so that the condensation with its corresponding high heat transfer coefficient can immediately start. As may be seen from the enthalpy – entropy (h-s) - diagram in Figure 2.33 the throttling process is isenthalpic, i.e. the enthalpy h of the flow before and after the throttle stays the same. However, a strong increase in the entropy may be observed, as the change of state is accompanied by large losses, i.e. irreversibilities. This illustrates once more that the supply of a thermal seawater desalination plant using a steam boiler represents thermodynamically a poor way of energy utilisation. As a result of the injection, the steam mass flow that condenses in the evaporator plant is not equal to the mass flow, which was produced by the steam boiler. However, by using a mass and energy balance a calculation can be made to obtain the steam, which enters the desalination plant: Mass balance

mHS ,out

mHS ,in minj

Energy balance mHS ,out hHS ,out

(2.95)

mHS ,in hHS ,in minj hinj

(2.96)

Due to the fact that the heating steam entering the evaporator or the brine heater has to be saturated steam and that the injected water is saturated liquid at the same temperature, the following is valid:

hHS ,out

hHS ,in

hV ,Tout

In order to obtain the mass flow of the heating steam we have to put these results in the above given mass-energy balances.

mHS ,out

mHS ,in

hHS ,in hTout hV ,Tout

(2.97)

The enthalpies for saturated steam and saturated liquid at Tout can be read out from the steam table [1]. Presuming that the mass flow, which is produced in the boiler and the enthalpy of the mass flow is given (usually by measuring the temperature and pressure) the heating steam mass flow can be directly calculated from Eq. (2.97). The fuel required for generating the heating steam for the desalination process can be determined using the efficiency of the boiler and the Lower Heating Value LHV of the fuel: B

mFu

Q LHVFu

(2.98)

116

Desalination 2nd Edition

Following the calculation of the pressure reduction station above, the heat which is transferred is given by:

Q

mHS ,out

hV ,Tout

(2.100)

That means the mass of fuel can be expressed as,

mHS ,out

mFu

B

hV ,Tout

LHVFu

(2.101)

It is now possible to define a modified Gained Output Ratio as the ratio of the fuel mass flow to the distillate produced as:

mD mFu

(2.102)

mD B LHVFu mHS ,out hV ,Tout

(2.103)

GORFu With Eq. 2.101 this gives:

GORFu

In section 3.1 the conventional GOR was defined as the ratio of the distillate mass flow to the heating steam mass flow.

GOR

mD mHS

By coupling of the two GOR´s together, we can see that the modified Gained Output Ratio is given as:

GORFu

GOR

B

LHVFu hV ,HS

(2.104)

In this definition, the GORFU can be applied as an efficiency suitable for all thermal processes, which are coupled to a steam boiler in singlepurpose mode. However, what is its practicability like? Let us continue with the example begun above using the following conditions for the boiler:

LHVFu

39, 400

kJ kg Fu

Heavy oil

Thermal Desalination Processes

117

Grid

Transformer Steam turbine

G ~

Electric power

2 1 Fuel

Desalination Dump condenser Steam boiler

Brine heater / evaporator 1stcell

High pressure pump

1

Pressure control valve

2

High pressure reducing station

Figure 2.34 Dual purpose plant/ Case 1: Back Pressure turbine.

B

0.75

hV ,HS

2, 338

kJ kg

As a result, we obtain:

GORFu

6.93

0.75 39, 400 2, 338

87, 68

The result makes it obvious why working with an efficiency defined in this way is so arduous: What do these “87.68” mean at all? Is it good or bad, high or low? An efficiency figure is clear if it is between 0 and 1. How practical is such a figure? It is immediately clear that the good old GOR always succeeds in defending its place for the assessment of energy utilisation in the literature when a direct coupling to the number of stages of the thermal plant is given. However, we would like to stick to this definition for the time being and see if this dimension is of use in comparing dual purpose plants with single purpose-plants. The simplest and most economical way of coupling a steam turbine power station and a thermal seawater desalination plant is the back pressure

118

Desalination 2nd Edition

1.

Condensation power station Q TTα Steam turbine

Exhaust

Pel

TC

Fuel Air Steam boiler

Condenser

high pressure pump 2.

steam generator Q Condenser Exhaust TC

Desalination

Fuel Air Steam boiler

high pressure pump 3.

back pressure turbine Q TTα

Condenser

Turbine

Exhaust

Fuel

TTω

Desalination

Air Steam boiler

High pressure pump

Figure 2.35 1) Condensation power station for electricity generation. 2) Steam generator as heat supply for a desalination plant. 3) Condensation power station with back pressure turbine.

Thermal Desalination Processes

119

turbine, as illustrated in Figure 2.34. In this the steam produced in the boiler firstly passed to a turbine, which produces electricity. This process should stop at a pressure/temperature level suitable for the requirements of the thermal seawater desalination plant, for example 120°C/2 bar for MSF plants. The seawater desalination plant takes the place of the power station condenser to all intents and purposes. The main disadvantage of this system is its lack of flexibility: if the turbine has to run at low load, then the seawater desalination plant also has to reduce its water production rate, which – depending on the type of plant – is not so simple. For this case and also for the case in which the turbine or the seawater desalination plant completely breaks down, a dump condenser and a by-pass around the turbine with a pressure reduction station must be planned. However, this creates extra costs and should be carefully considered during the design phase. In dual purpose-plants two “products” are generated from the fuel: electrical energy and heating steam. In using such a coupled production, the question may be asked as to how the total production costs may be apportioned across the two products. One possibility, which has the advantage of being justified by thermodynamics, uses the hypothetical approach of non-produced electricity. In these three systems are compared as represented in Figure 2.35: 1. A conventional condensation power station 2. A steam generator for the supply of the desalination plant 3. A power station with a back pressure turbine for supplying the desalination plant There are three relevant temperatures that have to be taken into consideration: 1. Inlet temperature of the turbine TTα 2. Outlet temperature of the turbine Tw 3. Condensation temperature TC The maximum work, which can be produced in an ideal cycle between two temperatures is given by the Carnot Factor, i.e. we can note the following for a condensation power station:

Work out C , PP

Heat in

TT TC TT

(2.105)

120

Desalination 2nd Edition

The work produced is pure energy, that means exergy ‘Ex’, so that Eq. 2.105 may also be formulated as follows:

ExT

TT TC Q TT

(2.106)

If a seawater desalination plant is connected to such a power station, then less electricity is produced in the turbine, i.e. a modified Carnot Factor is given as:

TT mod,Carnot

TT

(2.107)

TT

or

Ex mod,T

TT

TT TT

Q

(2.108)

The exergy, which may be ascribed to the seawater desalination, and which at the same time represents the loss of electricity produced, reads:

Ex HS

TT TC Q TT

C , HS

Q

(2.109)

The total fuel supplied to the condensation power station may be theoretically divided into a share which may be ascribed to the electricity generation and a share to the non-produced electricity, in other words the seawater desalination process.

mFu,total

mFu,T

mFu,DP

(2.110)

The chemical energy stored in a fuel is pure exergy, so that further formulae may be compiled:

mFu,total H Fuel

Ex PP

TT TC Q TT

(2.111)

Thermal Desalination Processes

mFu,DP H Fuel

TT TC Q TT

Ex HS

121

(2.112)

If these two equations are divided in themselves, and if Eq. 2.105 and 2.109 are taken into consideration the following results for the ratio of the two fuel flows:

mFu,DP mFu,total

Ex HS Ex PP

C , HS

(2.113)

C , PP

Now it is true that in all three cases (cf. Figure 2.35) the same amount of heat is produced in the boiler and that the same mass of fuel is required:

Q

mFu,total LHVFu

mHS

B

hV ,HS

(2.114)

If this equation is used to replace the total fuel flow in Eq. 2.113 we may obtain for the fuel flow of the desalination plant:

mFu,DP mHS

hV ,HS LHVFu

C , HS B

(2.115)

C , PP

It is helpful to introduce the total efficiency of a power station which comprises the efficiencies for the boiler, the turbine and the Carnot Cycle: PP

B

C , PP

(2.116)

T

Thus Eq. 2.115 becomes:

mFu,DP mHS

hV ,HS LHVFu

T

C , HS

(2.117)

PP

This equation has the advantage that the fuel mass flow tracable to the seawater desalination plant in dual purpose mode, can be calculated without knowing the relevant temperatures in the power station. It is sufficient to know the conditions of the desalination plant, the efficiency of the power station and the turbine and the type of fuel used. A Gained Output Ratio based on the fuel may be defined from this equation analogously to the procedure in single purpose mode (cf. Eq. 2.102):

GORFu,DP

mD mFu,DP

(2.118)

122

Desalination 2nd Edition

Using the conventional definition of the GOR for a thermal desalination plant, and incorporating Eq. 2.117, the following is finally obtained for dual purpose mode:

mD mHS

GOR GORFu,DP GORFu,DP

GOR

GOR

mHS mFu,DP

LHVFu hV ,HS

(2.119)

PP T

(2.120)

C , HS

If we compare this term with the result for the single purpose mode,

GORFu,SP

GOR

LHVFu hV ,HS

B

we see that the efficiency of the boiler was replaced by a combination of efficiencies. In the case of the single purpose mode an example was calculated above. In the following text a dual purpose plant will be calculated with the values used there. For this it is necessary to give the efficiencies and the relevant temperatures: Example: Efficiency of power plant

ηPP

0.35

Efficiency of the turbine

ηT

0.7

Heating steam temperature

TV,HS

68.3 0 C

Condensation temperature

TC

35 0C

Heat of evaporation

ΔhV,HS

2,385 kJ/kg

Net calorific value

LHVFu

39,400 kJ/kg

Gained output ratio

GOR

6.93

Efficiency of the turbine

ηB

0.75

First of all the Carnot efficiency to be ascribed to the desalination plant must be determined:

C , HS

TT TC TT

68.3 35 68.3 273.15

0.097

Thermal Desalination Processes

123

Using this and the equations given above, the result is:

GORFu,DP

6.93

39, 400 0.35 2338 0.7 0.097

602

In comparison with the case of a stand-alone steam boiler (single purpose):

GORFu,SP

6.93

0.75 39, 400 2338

87, 68

Therefore in this example a dual purpose plant has a GOR (related to the fuel demand), which is almost 7 times larger than that of a single purpose plant – in other words: the costs for the fuel are 7 times lower in dual purpose mode as in single purpose mode. Before further technical possibilities for the coupling of a power station with a desalination plant are investigated, a final remark on the efficiencies, in particular to the fuel-related GORFu should be allowed. The procedure detailed has shown that anyone working with these formulae is well advised to be clear on the conditions under which these formulae are valid. Both the classical GOR as well as the fuel-related GORFu can lead to quick and sound results when they are applied correctly. If you enter current values for the fuel and the efficiency rates (as in the above example, for instance), then both of the GORs can be relatively simply transferred between each other. The following rule of thumb may be utilised: Grid 1 Pressure control valve

3 Steam turbine

HP

G ~ LP

2 High pressure reducing station 3 Live steam valve

2

1 Fuel

Desalination Condenser Steam boiler Brine heater / evaporator 1st cell

High pressure pump

Figure 2.36 Dual purpose plant - Case 2: Controlled extraction steam turbine.

124

Desalination 2nd Edition

Rule of thumb for MSF plants (top brine temperature ≈ 120 °C) Single Purpose Steam generator

GORFu,SP Dual Purpose

15 GOR

Steam power station with back pressure turbine

GORFu,DP

50 GOR

Rule of thumb for MED and MED/TVC plants (top temperature ≈ 70 °C) Single Purpose Steam generator

GORFu,SP Dual Purpose

15 GOR

Steam power station with back pressure turbine

GORFu,DP

90 GOR

A general transfer to other possibilities for dual purpose operation, such as for example combined cycle or a block type power station, should be strictly avoided. The fuel demand, and its associated costs, should be determined on a case by case basis using mass and energy/exergy balances. 2 High pressure reducing station 3 Live steam valve 3 HP

4 Steam turbine

Steam turbine LP

4 Extraction control valve

2 Fuel

Desalination Condenser Steam boiler

Brine heater / evaporator 1st cell

High pressure pump

Figure 2.37 Dual purpose plant - Case 3: High pressure turbine + low pressure turbine.

Thermal Desalination Processes

125

Grid 1 Mechanical energy Compressor

Gas turbine

G ~ Electric power

Pressure control valve

1 Desalination

Air Fuel

Brine heater / st evaporator 1 cell

Combustion

Fuel EHR

Boiler feed water pump

Exhaust recovery boiler

Figure 2.38 Dual purpose plant - Case 4: Gas turbine + Exhaust Recovery Boiler.

Figure 2.36 shows the coupling of a desalination plant with a so-called extraction steam turbine. Here a part of the steam is taken out of the turbine at the pressure/temperature level (e.g. 2 bar/120 °C for MSF plants), which is necessary for the operation of the seawater desalination plant. The remaining steam continues to condense in the turbine and produces electricity. With the help of the valves 1, 2, and 3 which regulate the volume of the flow it is theoretically possible to set to any proportion of electricity to water production. The extraction steam turbine distinguishes itself from the back pressure turbine specifically through this flexibility. However, due to the increased amount of process control systems and the design of the turbine with an extraction port it is more costly to procure. The larger the power station and the larger the seawater desalination plant, the more varied the alternatives for coupling both, which cannot all be dealt with at this point. Figure 2.37 shows, as a further possibility, the constructional segmentation of the turbine into a high-pressure part and a low-pressure part. The desalination plant is situated exactly between these two parts because of the pressure/temperature level. Therefore behind the high-pressure turbine the steam is divided into two parts: one part flows to the desalination plant, the rest of the steam flows to the low-pressure turbine. A control of the ratio of electricity to water production is possible with the aid of the extraction control valve. Figure 2.38 shows the coupling of a desalination plant with a gas turbine. In this the exhaust gas from the gas turbine are passed to a steam generator which provides the heating steam for the thermal seawater

126

Desalination 2nd Edition

desalination plant. The operation mode at a constant load is extremely simple, as only one pressure control valve is required before the desalination plant. Uneconomic operation may be found on partial load, i.e. the gas turbine no longer delivers the volume of exhaust gas necessary for full load operation of the desalination plant. The loss in capacity can be partially compensated by the installation of an additional burner in the heat recovery boiler, whereby this layout then becomes more and more like a single purpose plant with a steam generator and all its associated negative characteristics. As gas turbine power stations are normally not base load power stations it should be checked exactly if and how they can be coupled to a seawater desalination plant. In principle there are no new aspects which have to be taken into consideration in the combination of a combined cycle with a seawater desalination plant: the exact design of a dual purpose plant can in the end be seen in the context of the electricity and water demand of a particular region or even of the country in question. Above all the aim of this chapter is to make clear that we should prefer the dual purpose mode in thermal seawater desalination in order to optimise the consumption of the chemical energy stored in the fuel.

2.3.3 Specific Primary Energy Consumption When we speak of primary energy we mean the chemical energy stored in the fuel which can be regarded as pure exergy:

Ex

mFu H Fuel

(2.121)

In order to determine the primary energy demand a calculation must therefore be made backwards from the energy demand of the plants to the exergy and with this to the fuel. In doing this a difference must be made between the thermal and electrical energy to be supplied to the plants. The Assuming typical figures for desalination plants was already calculated as follows in section 3.2 (cf. Eq. 2.101 and Eq. 2.117): Single purpose

Dual purpose

mFu,SP

mFu,DP

mHS

hV ,HS B

LHVFu

mHS

hV ,HS LHVFu

T

C , HS PP

Thermal Desalination Processes

127

Using Eq. 2.121 the exergy flow may be determined as: Single purpose

Ex SP ,th

mFu,SP LHVFu

mHS

1

hV ,HS

(2.122)

B

Dual purpose

Ex DP ,th

mFu,DP LHVFu

mHS

hV ,HS

T

C , HS PP

(2.123)

Although the heat represents the major part of the primary energy cost for the thermal plants, the demand of electrical energy for pumps and auxiliary equipment must not be forgotten. Without going into the individual items in greater detail, at this point an approximated value for the different plants should be used: MSF plants (Brine-recycle mode)

MED plants

Pel , MSF

Pel , MED

mD

2. 5

kWh m‡

2. 0

mD

kWh m‡

The electrical energy is produced in a power station. Therefore, the fuel consumption can be directly derived via the total efficiency of the power station: PP

mFu

Pel LHVFu

(2.124)

This results in the following expression for the exergy demand for provision of electrical energy: MSF plants (Brine-recycle mode)

Ex MSF ,el

mFu, MSF LHVFu

MED plants

Pel , MSF

Ex MED ,el

mFu, MED LHVFu

PP

Pel , MED PP

The heat demand of an MSF plant and an MED plant can be expressed in a formula as follows, as shown in section 2.3: MSF (cf. Eq. 2.47):

QH , MSF mD

hV ,Tm N

1 N

TTTD

TBPE TO

TLosses

128

Desalination 2nd Edition

hV ,TV ,m

QH MED (cf. Eq. 2.27): m D

N

CF c ( TStage CF 1 P

TTTD

TBPE )

The heat is provided by the condensation of saturated steam, so that the following terms apply:

QH , MSF

mHS , MSF

hV ,HS

(2.125)

QH , MED

mHS , MED

hV ,HS

(2.126)

If you now combine Eq. 2.125, Eq. 3.38 with Eq. 2.47, Eq. 2.67 respectively (the single mode may be treated in an analogous manner which does not happen here so that we can retain a clear overview), then, in each case, you obtain an equation for the specific exergy requirement of an MSF or an MED plant: MSF/dual purpose

Ex DP ,th, MSF

hV ,Tm

mD

1 N

N

TTTD

TBPE T0

TLosses

T

C , HS PP

(2.127) MED/dual purpose

Ex DP ,th, MED mD

hV ,Tm

CF c ( TStage CF 1 P

N

TTTD

TBPE )

T

C , HS PP

(2.128) In order to arrive at the total specific exergy demand, and thus the specific primary energy demand of the plants, the share for the electrical energy must be added to the heat. Using approximated values this finally gives: MSF/dual purpose

Ex DP , MSF mD

hV ,Tm N

1 N

TTTD

TBPE T0

TLosses

T

C , HS PP

Pel , MSF PP

mD (2.129)

Thermal Desalination Processes

129

Table 2.1 Typical values of thermal and RO desalination plants as a basis for Figure 2.39. Ex

Primary Energy Consumption

[kW]

N

Number of Stages

[-]

m ˙D

Mass Flow Distillate resp. Permeate

[to/h]

ΔTTTD

Terminal Temperature Difference

1.5 K

ΔTBPE

Boiling Point Elevation

0.7 K

ΔTLOSSES

Non-Equilibrium Losses

0.6 K

ΔT0

Overall Temperature Difference

80 K

TT

Temperature of Heating Steam

393 K

TMIN

Theoretical Condensing Temperature

313 K

ΔhV

Heat of Evaporation @ 393 Degree K

2,200 kJ/kg

cp

Specific Heat Capacity of Seawater

4.0 kJ/kg K

ηpp

Efficiency Power Plant

0.4

ηT

Efficiency Steam Turbine

0.9

ηp

Efficiency RO High Pressure Pump

0.8

ηERT

Efficiency RO Energy Recovery Turbine

0.8

ΔpRO

Driving Pressure Difference

30 bar

ΔpLOSSES

Pressure Losses Due to Friction

4 bar

b

Osmotic Pressure Coefficient

8 bar/%TDS

wF

Feed Concentration

4.2 %TDS

CFRO

RO Concentration Factor

1.5

CFMED

MED Concentration Factor

2.0

ρ

Density Of Saline Water

1,041 kg/m3

PAUX RO/m ˙D

Specific Energy Demand for Auxiliaries RO

0.5 kWh/to

PAUX MSF/m ˙D

Specific Energy Demand for Auxiliaries MSF

2.0 kWh/to

˙D PAUX MED/m

Specific Energy Demand for Auxiliaries MED

2.0 kWh/to

MED/dual purpose

Ex DP , MED mD

hV ,Tm N

CF c CF 1 P

TStage

TTTD

TBPE

T

C , HS PP

Pel , MED PP

mD (2.130)

Desalination 2nd Edition

130

Specific primary energy consumption Ex/mD [kWh/t]

100

80

Multi stage flash (MSF)

60

Multiple effect (ME)

40

Reverse osmosis (RO)

20

0 0

10

20

30

40

50

60

Number of stages N [ - ]

Figure 2.39 Specific exergy and primary energy consumption vs. number of stages. (MSF and MED: Dual- purpose mode).

As reverse osmosis only needs electrical energy the determination of the exergy demand is relatively simple.

Pel ,RO mD

1

pRO

pLOSS b w F CF 1 1 P 1 CF

pRO b w F CF CF 1

ER

(2.131) Assuming typical figures for desalination plants according to the stateof-the-art (see Table 2.1) a graph may be drawn, which shows the specific exergy demand and the specific primary energy consumption respectively as a function of the number of stages N. In Figure 2.39 three curves may be identified: 1. MSF plant/dual purpose mode 2. MED plant/dual purpose mode 3. RO plant

→ → →

Eq. 2.128 Eq. 2.129 Eq. 2.130

The following may be observed: The primary energy demand of reverse osmosis amounts to approximately 15 kWh/m3 independently of the number of stages.

Thermal Desalination Processes

131

As was to be expected there is a strong dependency on the number of stages for the thermal plants. MSF plants always require more primary energy than MED plants. However, a state-of-the-art MSF plant (approx. 20 stages, GOR ≈ 8) requires about 30 kWh/m3; state-of-the-art MED plants (approx. 10 stages, GOR ≈ 8) require around 35 kWh/m3. Only above a number of stages of N > 40 do MED plants catch up with RO plants, MSF plants only reach values of 35 kWh/m3 at the most, even with a high number of stages.

2.4 Recent Developments in Thermal Desalination Processes 2.4.1 Hybrid Plants We have learned in the last section that dual-purpose thermal desalination plants lead to an essential reduction of the energy cost of water production. There is another option to reduce the energy cost, which is gaining popularity these days – namely hybrid plants [5]. “Hybrid” means that two different types of desalination plants are coupled together. Most often used is a combination of thermal desalination with membrane desalination processes.

2.4.1.1 Multi-Stage Flash with Reverse Osmosis (MSF-RO) In this particular hybrid plant, Multi-Stage-Flash (MSF) desalination plant is combined with Reverse Osmosis (RO) plant. The MSF process draws heating steam from the power plant, which is waste steam and uses this heat in the steam to preheat seawater as required in the MSF unit. The RO plant uses electricity generated by the power station and operates during periods of reduced power demand, thus optimizing the overall efficiency of the entire operation. One of the largest hybrid seawater desalination plant combined with a gas turbine is located in Ras Al-Khair in Saudi Arabia. The desalination plant has a total capacity of 228 MIGD and consists of 8 MSF units and 17 RO units. 70% of the total amount of fresh water is produced by Multi-Stage-Flash Distillation plant and 30% by Reverse Osmosis. The plant is operated by the Saline Water Conversion Corporation (SWCC). The Combined Cycle Gas Turbine power station is producing 2,400 MW electricity.

Desalination 2nd Edition

132

2.4.1.2 Multi-Effect Distillation with Reverse Osmosis (MED-RO) The largest seawater hybrid desalination plant combining Multi-Effect Distillation (MED) and Reverse Osmosis is located in the United Arab Emirates and is called Fujairah 2. This plant has actively operated since May 2011 with net capacity of 100 MIGD produced by MED and 30 MIGD produced by RO. The gas turbines are producing 2,000 MW of electric power [6].

2.4.2

Expanding the Scope of Hybrid Thermal Desalination

Option MED-MSF There are even some suggestions to combine MSF and MED plants. In this regard, some theories and researches are presented, which will be beneficial for water generation at low cost and with proper utilization of energy. It will be interesting to see if this turns out to be one of the practical solutions for better future [5]. Employing Filtration to Thermal Processes By looking at the advantages offered by membrane separation technologies, researchers are very keen to propose a vision for the future of coupling membrane technology and thermal seawater desalination plants. The coupling of MSF or MED plants with a Reverse Osmosis is already considered and employed in many plants as it has shown potential to be more economical than stand-alone thermal/membrane desalination plants. This hybrid plant has the advantage that no conventional pre-treatment

Nano filtration (NF) unit

Pre-treatment Seawater in

NF product

Seawater in

Steam MSF Condensate

Seawater out MSF product

Figure 2.40 Block diagram of an Nanofitration (NF) - Multi-Stage-Flash (MSF) hybrid plant.

Thermal Desalination Processes Pre-treatment

133

Nano-filtration (NF) unit

Seawater intake NF product

NF reject

Combined product

RO product MSF product

RO

Steam

Seawater in MSF

Seawater out

Condesate RO reject

Figure 2.41 Block diagram of Tri-hybrid plant comprising Nanofiltration (NF) – Reverse Osmosis (RO) and Multi-Stage- Flash (MSF) [8].

of the seawater would be required. In addition to RO, Ultrafiltration (UF) and Nanofiltration (NF) could also be merged with thermal desalination plants. Especially Nanofitration could significantly reduce the cost of the processes by eliminating the risk of scaling and fouling as it will allow an increase in the top temperature, which results in a significant increase in output and the performance ratio. Different combinations of materials for are being tested and accordingly developed to employ this efficiently [7]. To avoid scale formation in MSF, an innovative concept using NF membrane for pre-treatment of feed to MSF desalination plant is introduced. Tests are being performed to check the feasibility of the process. Below in Figure 2.40 is the block diagram of such process. This shows the seawater with some necessary pre-treatment passes through NF process. Then the output of this is transferred to MSF as a feed. After MSF desalinated water is transfer for storage and rest brine to sea as a reject. The latest proposal is to combine three existing technologies together as shown in Figure 2.41. The task of the Nanofiltration is to remove the hardness, mainly Calcium and Magnesium, in order to reduce the risk of scaling. The MSF plant takes the reject of the RO plant. Some research papers have already shown the economical side of this hybrid plant when compared to individual plants for same amount of water generation. NF pre-treatment enables high temperature operation of MED without scale formation, which means increased production of water with reduction in production costs [8].

134

Desalination 2nd Edition

To check the feasibility and study the benefits of this fully integrated NF-RO-MED tri-hybrid plant, a test plant was erected and commissioned at Al Jubail by the Saline Water Desalination Research Institute (SWDRI) in 2008.

2.5 Future Prospects 2.5.1 General Remarks What does the future of seawater desalination look like? Will there be any alternatives to what is used in the state-of-the-art? What will be a significant source of energy for thermal desalination processes? Research nowadays is on new technologies and simplified design of existing technologies. Following are some of the topics, which will play an important role in the future and are to be discussed later: Minimize and control scaling and fouling Different materials of construction Optimization of process design Control system to optimize consumables consumption Hybrid desalination plants to reduce cost and to increase volumes of desalinated water Utilization of different energy sources other than conventional This can further be categorized as optimization of existing process design in order to improve efficiencies and economical operation and as proposal of new technologies in order to fulfil energy demands in the future.

2.5.2 Optimization of Existing Process Design This is crucial as researchers already have an idea about the operational conditions and its flaws. Which areas are to be rectified so that the system will give high-efficiency output? This includes the substitution of costly materials by cheap ones but without compromising on strength and securities.

2.5.2.1

Material of Construction

Low-temperature evaporators, able to operate below 70 °C with waste heat, could be developed as a counterpart to the high-temperature evaporator

Thermal Desalination Processes

135

plants. The installation of a higher number of stages assumes small temperature differences for the heat transfer, which would directly lead to the need for larger heat transfer areas. The evaporators which would have to be developed must therefore consist of cheap materials, which are corrosionresistant and guarantee a high heat transfer coefficient. Thin plastic sheets can well meet these demands; however, it has not yet been definitely clarified as to how the wetting of the sheets could be made without scaling occurring. Plastic tubes are a possible alternative, in which case the question may immediately come up about the mechanical stability and the wall thickness. Scale formation is a result of saline water and hot conditions. A simple surface painting or surface coating as corrosion protection is not sufficient. The use of corrosion-resistant alloys (CRA) are needed to prevent corrosion effects, which degrades the performance of material. For the most corrosion-prone parts Titanium is used, whereas copper alloys are usually chosen for less critical heat exchangers. For high fluid velocity application Titanium and stainless steel can be used because of high erosion resistance [9]. Efforts are being made to substitute existing metal components like heating surfaces, structural parts and tubes with highly filled polymer having similar or enhanced strength and heat transfer properties. High thermal capacity materials can be used here as a filler, which is advantageous. This will reduce the weight of the system part, thereby possibly reducing component cost. Use of this material will not only offer a reduction in investment costs but also in operating costs due to less use of anti-scaling substances and reduced cleaning and maintenance purposes. Also the operating temperature of thermal desalination process (MED) could be increased since this helps in scaling prevention. The ultimate outcome of this will be improved Gained Output Ratio. These thermal and structural properties of the materials need further research and analysis to verify their feasibility [10].

2.5.2.2 Increasing Water Velocity The velocity of brine in the tube could play an important role in desalination plants. Low velocity of water allows suspended salt particles to settle down, which prevents adequate water distribution across the tube bundle. The use of anti-scalant and acid can be used to keep the brine velocity to a desired level. The permissible water velocity is dependent on the tube material. Cupronickel materials are known to be subjected to erosion under high velocity while titanium and stainless steels are quite resistant. High vapour release rate at the brine surface results in increased liquid entrainment, which is also strongly influenced by brine velocity. It is also to

136

Desalination 2nd Edition

be recognized that demister height above the brine surface greatly affects performance, product quality and production longevity. This certainly will play an important role in improving the performance of the desalination plant.

2.5.2.3 Heat Transfer Enhancement by Using Corrugated Oval Tubes According to the experiments and analysis performed, vertical fluted tubes are considered capable of providing superior heat transfer coefficients. Though this type of tube is little costlier than a regular one ultimately it serves as economical in operation. Such design of tubes can be used in thermal desalination plants [11].

2.5.2.4 Increasing the Top Operation Temperature to 85 °C By increasing the top brine temperature from 70 °C to 85 °C would be considered as good as increasing the number of the effects, thus increasing the GOR and thereby reducing the water-cost. But to employ these various other changes should be taken into consideration like material of heat transfer surface etc. [11].

2.5.2.5

Increasing Number of Stages

A higher number of stages mean in any case a reduction in the thermal energy consumption of the plant but also the construction of a larger number of stages with the relevant evaporators and preheaters, which also have to operate at higher pressures. Therefore, an increase in the number of stages is only worthwhile where cheap materials can be used for the heat exchanger and a cheap design can be made. In addition, the processrelated thermodynamic and hydrodynamic losses increase with the number of stages, so that the energy saving from a certain number of stages is only marginal.

2.5.2.6 Modifications in MED-TVC MED is the oldest thermal desalination process but in previous decades emphasis was given on MSF technology as MED is not economic when compared to MSF. Due to this MED-TVC is introduced, which is developing rapidly and most preferred all over as it is more energy efficient. Researchers are proposing to change the position of thermal vapour

Thermal Desalination Processes

137

compressor in order to improve its performance and reduce energy consumption. Software simulation is being employed to know the significance of this proposal by different researchers. There are some other proposals too which suggest replacing the thermally driven compressor by a large centrifugal compressor. In a thermally driven compressor the losses are significant, which reduces the efficiency, thereby increasing energy consumption. To avoid these losses a large centrifugal compressor driven by an auxiliary steam turbine utilizing extracted steam is suggested. Centrifugal compressor and auxiliary steam turbine has comparatively higher efficiency, which may result in lowering energy consumption, thereby reducing the production cost of water [11].

References 1. J. Gebel and S. Yüce, An Engineer’s Guide to Desalination, VGB Powertech Service GmbH, Essen. (2008) ISBN 978-3-86875-000-3. 2. B.P. Robert, Steam Jet Ejections for the Process Industries, McGraw-Hill, Inc., (1994) ISBN 0-07-050618-3. 3. VDI, Wärmeatlas Berechnungsblätter für den Wärmeübergang, VDI-Verlag, Düsseldorf 5, Auflage, (1988) ISBN 3-18-400760-X. 4. H.E. Hömig, Seawater and Seawater Desalination, Vulkan-Verlag, Essen 1978, ISBN3-8027-243-0 5. Osman A. Hamed, Overview of hybrid desalination systems - Current status and future prospects International Conference on Water Resources and Arid Environment, Riyadh, Saudi Arabia, 5–8 December 2004. 6. Leon Awerbuch, New advances and future needs in thermal-membrane hybrids for water desalination, World Congress/Perth Convention and Exhibition Centre (PCEC), Perth, Western Australia September 4–9, 2011, REF: IDAWC/PER11-378 (987) 7. Mabrouk Methnani, Seok Ho Choi, Altaf Hussain, Young Jin Baek, Technoeconomics of Hybrid Desalination Systems, 2007, IDA World CongressMaspalomas, Gran Canaria –Spain October 21–26, 2007, REF: IDAWC/ MP07-036 (1838). 8. Osman A. Hamed, Takaji Akiya, Hiroshi Miyamura, Toru Kannari and Koji Harada, Development of 7 MIGD MED-TVC distiller within the context of tri-hybrid NF/RO/MED configuration World Congress/Perth Convention and Exhibition Centre (PCEC), Perth, Western Australia September 4–9, 2011; REF: IDAWC/PER11 – 184 9. Haydée V. Richaud-Minier, Pascal Gérard, Hervé Marchebois, Use of a New Alternative Solution for MED and MSF Heat Exchanger Tubing IDA World Congress-Maspalomas, Gran Canaria –Spain October 21–26, 2007, REF: IDAWC/MP07-021

138

Desalination 2nd Edition

10. Thomas Orth, Dirk Moses, Sebastian Ossadnik, Patrick Mueller, Norbert Scherer, Alex Dantziguian, A new class of material for thermal desalination: High performance polymer compounds, The International Desalination Association World Congress on Desalination and Water Reuse 2015/San Diego, CA, USA, 2015, REF: IDAWC15- Orth_51460(4319) 11. A. Ophir, A. Gendel, Latest Developments in MED and MVC Thermal Desalination Processes IDA World Congress-Maspalomas, Gran Canaria – Spain October 21–26, 2007, REF: IDAWC/MP07-067(1862)

3 Basic Terms and Definitions Mark Wilf European Desalination Society at the Genoa University, Italy and San Diego, CA, USA

Abstract This chapter defines basic terms used in conjunction with RO systems. Also see Chapter 9 for additional information as to how these parameters affect the performance of an RO system.

3.1 Reverse Osmosis System Flow Rating An RO system is rated based on product flow rate. An 800-gpm RO would yield 800 gpm of permeate. The influent and reject flows are typically not indicated except in the design details (they are usually calculated knowing the product flow rate and the percent recovery). In some cases, the actual design permeate flow rate of the RO system may differ from the “name plate” flow rating. In most of these situations, the RO system is de-rated by design due to a poor feed water source or as a natural result of low feed water temperature.

Jane Kucera (ed.) Desalination 2nd Edition, (139–162) © 2019 Scrivener Publishing LLC

140

Desalination 2nd Edition

3.2 Recovery Recovery (sometime referred to as “conversion”) is a term used to describe what volume percentage of influent water is “recovered” as permeate. Generally, RO system recoveries range from about 50% to 85%, with the majority of systems designed for 75% recovery. (Individual spiral wound membrane module recoveries vary from about 10% to 15%— see Chapter 4.3). A system recovery of 75% means that for every 100 gpm influent, 75 gpm will become permeate and 25 gpm will be retained as concentrate. Recovery is calculated using the following equation: % Recovery = (permeate flow/feed flow) * 100

(3.1)

At 75% recovery, the concentrate volume is one-fourth that of the influent volume. If it were assumed that the membrane retains all the dissolved solids, they would be contained in one-fourth of the volume of influent water. Hence, the concentration of retained dissolved solids would be four times that of the influent stream (since not all dissolved solids are retained by the membrane, this becomes only an approximation). This is called the “concentration factor.” At 50% recovery, the concentrate volume would be one-half that of the influent water. In this case, the dissolved solids would be concentrated by a factor of two, so the concentration factor would be 2. Table 3.1 shows the concentration factor as a function of recovery. Understanding the reject concentration is important as the concentrate side of the membrane is the area where fouling and scaling occur (see Chapters 3.6 and 3.7 respectively). Higher recovery results in the need to dispose of less reject water. However, higher recovery also results in lower-purity permeate.

Table 3.1 Concentration factor as a function of recovery. Recovery (%)

Concentration factor

50

2

66

~3

75

4

80

5

83

6

87.5

8

Basic Terms and Definitions Instantaneous permeate concentration

2 ppm

4

141

8

98% Rejection membrane

Feed / Concentrate 100 ppm concentration recovery 0%

200

400

50%

75%

Figure 3.1 Concentrate and instantaneous permeate concentration as functions of recovery.

Consider the example shown in Figure 3.1. At the influent end of the membrane, the influent concentration is 100 ppm, while the recovery is 0%, and the membrane passes 2% total dissolved solids (TDS) (see Chapter 3.3). The permeate right at this spot would be about 2 ppm. As the influent water passes across more and more membrane area, more water is recovered. At 50% recovery, the concentration factor is 2, so the influent water now has a concentration of about 200 ppm. The permeate water at this point would now have a concentration of 4 ppm. At 75% recovery, the concentration factor is 4, so the influent water now has a concentration of about 400 ppm. The permeate water at this point would have a concentration of 8 ppm. Hence, higher recovery results in lower product purity. The designer of the RO system selects the recovery for the system; it is not a property of the membrane. The designer must consider the trade off between higher recovery resulting in less concentrate water to dispose of but also lower permeate purity. In practice, the recovery of the RO system is adjusted using the flow control valve located on the RO concentrate stream (see Figure 2.6). Throttling the valve will result in higher operating pressure, which forces more water through the membrane as opposed to down along the feed/concentrate side of the membrane, and results in higher recovery. The recovery of an RO system is fixed by the designer. Exceeding the design recovery can result in accelerated fouling and scaling of the membranes, because less water is available to scour the membrane on the concentrate side. Falling below the design recovery will not adversely impact membrane fouling or scaling, but will result in higher volumes of wastewater from the RO system.

142

3.3

Desalination 2nd Edition

Rejection

Rejection is a term used to describe what percentage of an influent species a membrane retains. For example, 98% rejection of silica means that the membrane will retain 98% of the influent silica. It also means that 2% of influent silica will pass through the membrane into the permeate (known as “salt passage”). Rejection of a given species is calculated using the following equation: % Rejection = [(Cf – Cp)/ Cf ] * 100

(3.2)

where: Cf = influent concentration of a specific component Cp = permeate concentration of a specific component Note that for exact calculation, the average feed concentration that takes in account both the feed and concentrate concentration rather than just the feed concentration at a single point in time should be used. Salt passage is essentially the opposite of rejection: % Salt Passage = 100 -% Rejection

(3.3)

% Salt Passage = (Cp/Cf ) * 100

(3.4)

Sometimes, it is easier to consider membrane performance in terms of what passes through the membrane than what is retained by the membrane. Rejection is a property of the specific feed water component and the membrane of interest. Table 3.2 lists the general rejection ability of the most common polyamide composite RO membranes. Note that ionic charge of the component of interest plays a key role its rejection by an RO membrane; the rejection of multivalent ions is generally greater than for mono-valent ions. In addition to the ionic charge, rejection of a particular species is also based on the following characteristics:[1] Degree of dissociation: in general, the greater the dissociation, the greater the rejection, for example, weak acids are rejected better at higher pH. Molecular weight: in general, the greater the molecular weight, the greater the rejection, for example, the rejection of calcium is marginally better than the rejection of magnesium. Polarity: in general, the greater the polarity, the lower the rejection, for example, organics are rejected better than water.

Basic Terms and Definitions Table 3.2 General rejection capabilities of most polyamide composite membranes at room temperature. Species

Rejection (%)

Sodium

92–98

Calcium

93 – 99+

Magnesium

93–98

Potassium

92–96

Iron

96–98

Manganese

96–98

Aluminum

96–98

Ammonium*

80–90

Copper

96–99

Nickel

96–99

Zinc

96–98

Silver

93–96

Mercury

94–97

Hardness

93–99

Chloride

92–98

Bicarbonate

96–99

Sulfate

96–99+

Fluoride

92–95

Silicate

92–95

Phosphate

96–98

Bromide

90–95

Borate

30–50

Chromate

85–95

Cyanide

90–99+

* below pH 7.5. Above this pH, a greater percentange of the ammonia exists as a gas which is not rejected by RO membranes.

Degree of hydration: in general, the greater the degree of hydration, the greater the rejection, for example, chloride is rejecter better than nitrate. Degree of molecular branching: in general, the more branching, the greater the rejection, for example, isopropanol is rejected better than normal propanol.

143

144

Desalination 2nd Edition

The rejection of gases is 0%, meaning that the concentration in the permeate stream will be the same as it is in the influent and concentrate streams. Gases that are not rejected include free chlorine that may used to disinfect RO feed water through the pretreatment system (see Chapter 8.2) and carbon dioxide. RO systems operating at near neutral pH will have some carbon dioxide in the feed water. Since gases are not rejected by RO membranes, the permeate and concentrate streams will also contain carbon dioxide. If the permeate is sent to ion exchange demineralization or electrodeionization after the RO, the carbon dioxide will use sites on the anion resin so that other anions are not well removed. In these cases, caustic soda (NaOH) is sometimes added to the RO feed water. This raises the pH and converts the carbon dioxide, which is not rejected by the RO membrane, to bicarbonate, which is rejected by the RO membrane. Caustic addition is recommended after sodium softening, which removes hardness (calcium, magnesium, barium, and strontium). Without softening, hardness in the feed water would saturate at the higher pH following caustic addition and scale the membranes. Caustic is also sometimes added between passes in a two-pass RO system (see Chapter 5.3); the firstpass RO removes the hardness while the effluent from the second pass is relatively free of carbon dioxide following caustic addition to the second pass feed. Because carbon dioxide passes through RO membranes, the pH of the permeate is lower than the pH of the feed stream for feed water with a pH lower than about 9. Any carbon dioxide in the feed will pass through the membrane while any bicarbonate will not. This changes the ratio of carbon dioxide to bicarbonate in both the permeate and the concentrate, with the permeate having a higher ratio of carbon dioxide to bicarbonate than the feed and the concentrate having a higher ratio of bicarbonate to carbon dioxide than the feed. Hence, the pH of the permeate will be lower than the feed, while the pH of the concentrate will be higher than the feed. Another gas that is not rejected by RO membranes is ammonia. Ammonia is a consideration when treating wastewaters as well as feed water that has been treated with chloramine. Figure 3.2 shows the relative concentrations of ammonia gas and ammonium ion as a function of pH. At a pH of approximately 9.3, half of the ammonia species is present as ammonia gas and half as ammonium ion. The gas is not rejected, while the ion has a rejection of upwards of 80% (see Table 3.1), making the overall rejection of ammonia typically less than 50%. To achieve a relatively high rejection of ammonia, the pH of the water to the RO membranes should be less than 7–7.5, as shown in Figure 3.2. Note that ammonia gas is known to swell polyamide membranes, which causes the rejection of salts to be

Basic Terms and Definitions

145

Ammonia gas & Ammonium ion equilibrium (based on C and low ionic strength solution) 110 100 NH4+

Relative concentration, %

90 80 70 60 50 40 30 20 10

NH3

0 -10 0.0

1.0

2.0

3.0

4.0

5.0

6.0

7.0

8.0

9.0 10.0 11.0 12.0 13.0 14.0

pH at 25 C

Figure 3.2 Concentration of ammonia gas and ammonium ion as functions of pH.

reduced. Salt passage can double when the membranes are exposed to free ammonia. However, this is a reversible condition, and once the free ammonia is removed, typically by reducing the pH of the water, the rejection of the RO membranes will return to normal. Membrane systems operating on city water treated with chloramine, particularly when breakpoint chloramination is occurring, can expect to see an increase in salt passage, should the pH be greater than about 8 (which is common for municipal water sources). Thus, when a system operating on city water experiences a sudden decrease in permeate quality, city workers should be consulted to determine if they are currently using chloramine.

3.4 Flux Flux is defined as the volumetric flow rate of a fluid through a given area. In the case of RO, the fluid is water and the area is that of the membrane. In the language of RO, flux is expressed as gallons of water per square foot of membrane area per day, (gfd). The flux of water through an RO membrane is proportional to the net pressure driving force applied to the water (see Chapter 4.1 for a discussion on transport models): J = K (ΔP – ΔΠ)

(3.5)

146

Desalination 2nd Edition

where: J = water flux K = water transport coefficient = permeability/thickness of the membrane active layer ΔP = pressure difference across the membrane ΔΠ = osmotic pressure difference across the membrane Note that the water transport coefficient is unique to a given membrane and is not a constant; it varies directly with temperature. The coefficient for some newer polyamide membranes also varies with pH. The designer of the RO system chooses the flux rate; it is not a property of the membrane. In general, the flux that an RO system is designed for should be a function of the influent water quality. This is because higher flux results in more rapid fouling of the membranes. So, the lower the influent water quality, the lower the operating flux of the RO system should be. Table 3.3 shows the recommended flux as a function of influent water source (which is an indirect measure of the water quality) and silt density index (SDI), which is a measure of the tendency of water to foul a membrane (See Chapter 3.9). When in doubt, a default flux of 14 gfd is usually recommended. Specific flux is sometimes discussed in comparing the performance of one type of membrane with another. Since not all membranes are tested at the same pressure. Specific flux is approximated by taking the overall system flux and dividing by the applied driving pressure: Specific Flux = Flux/Applied Pressure

(3.6)

Table 3.3 Recommended flux as a function of influent water source. Feed water source

SDI

Recommended flux, gfd *

RO Permeate

Pb2+ 88.7% (1.20 Å)> Ni2+ 86.1% (0.72 Å) > Zn2+ 81.4% (0.74 Å). This observation is in agreement with the experimental results of Gabelich et al. who reported that the ion selectivity of carbon aerogel electrodes was based on ionic hydrated radius [39]. Figure 12.5 shows the variation of conductivity and the metal concentrations in the treated water throughout the testing period. The electrical resistance setpoint of these experiments was 500 ohms in the experiments, the treatment time increased gradually from 25 minutes in the beginning of the testing to 30 minutes at the end of the experiments. The conductivity of the effluent increased slightly from 13 μS/cm to 22 μS/cm at the end of the testing. There are no drinking water standards set by the USEPA for nickel in water. Over the course of 158 runs, the nickel concentration in the treated water increased from 0.074 mg/L in the first few runs to between 2–5 mg/L for the Runs #50 to #158. CDI achieved 89.0±7.3% removal of nickel with the feed concentration of 18.74 mg/L.

Desalination 2nd Edition

512

10 Ni Concentration (mg/L)

Conductivity (uS/cm)

50 45 40 35 30 25 20 15 10 5 0

1

0.1

0.01 0

50

100

0

150

50

100 Runs

150

Runs 1

10 Cu Concentration (mg/L)

Fe Concentration (mg/L)

Primary MCL 0.3 mg/L 0.1

0.01

0.001 50

100 Runs

0.1

0

150

50

100

150

Runs 10

1

0.1 USEPA Action Level 0.015

Zn Concentration (mg/L)

10 Pb Concentration (mg/L)

Secondary MCL 1.0 mg/L

0.01 0

Primary MCL 5 mg/L

1

0.1

0.01

0.01 0

50

100 Runs

0

150

50

100 Runs

150

1 U-238 Concentration (mg/L)

10 Cr Concentration (mg/L)

Primary MCL 1.3 1

1 Primary MCL 0.1 0.1

0.01

0.001

0.1 Primary MCL 0.03 mg/L

0.01

0.001 0

50

100 Runs

150

0

50

100

150

Runs

Figure 12.5 Variation of conductivity and metal concentrations in the treated electrolyte solutions throughout the 158 runs (setpoint: 500 ohms, corresponding to 25–30 minutes contact time).

Electrosorption of Heavy Metals

513

Within 150 runs of the testing, the chromium concentrations decreased from 11.15 mg/L in the feed solution to below the MCL level of 0.1 mg/L regulated in the USEPA Primary Drinking Water Standards. At the end of the testing, the chromium concentration reached close to the MCL. CDI achieved 99.6±0.3% of chromium removal throughout the experiments. During the first 60 cycles, the copper concentrations were reduced from 15.56 mg/L in the feed solution to below the USEPA Drinking Water Secondary Standards of 1.0 mg/L in the treated water. After 75 runs, the copper concentrations in the effluent exceeded the Primary MCL of 1.3 mg/L. The average removal rate of copper was 94.8±2.7% throughout the 158 runs. In order to meet the desired removal of contaminant from the water, increase of the retention time is required. An alternative option is to have a number of CDI cells arranged in series using the end cells to meet the drinking water standards and the first few cells for recovery of metals from water. Throughout the testing, the iron concentrations in the effluent were below 0.04 mg/L, much lower than the Primary MCL of 0.3 mg/L. With the feed concentration of 16.09 mg/L, the percent removal of iron by CDI was 99.9±0.06% over the158 runs. USEPA set a lead concentration of zero in drinking water standards and an action level of 0.015 mg/L. The lead concentration in the treated water was between 0.4 and 2 mg/L exceeding the drinking water standards. With a feed concentration of 19.03 mg/L, CDI achieved 94.0±3.3% removal of lead during the testing. The zinc concentrations in the effluent increased with the treatment runs from the beginning of 0.2–0.3 mg/L to between 1–4 mg/L in the later runs. The percent removal of zinc by CDI was 85.2±9.0% throughout the testing and the zinc concentration was reduced from 12.09 mg/L to below the Primary MCL of 5 mg/L. During the experiments the uranium concentrations were reduced from 1.85 mg/L in the feed solution to the range of 0.005 to 0.043 mg/L in the treated water, achieving 98.9±0.6% removal. However, approximately 20% of the effluent samples slightly exceeded the Primary uranium MCL of 0.03 mg/L. Table 12.4 lists the metal concentrations in the samples collected from the anode and cathode drains. Because metals are positively charged ions and attracted by negative electrode, the metals moving towards the cathode adsorbed on the carbon materials and were removed from water solution. Thus the metals concentrations in the drains from the cathode chamber were low. However, nickel and zinc were found at concentrations

514

Desalination 2nd Edition

Table 12.4 Metal concentrations measured in the electrode drain samples. Concentration in drains from anode chamber (mg/L)

Concentration in drains from cathode chamber (mg/L)

Cr

0.71

0.002

Cu

1.13

0.004

Fe

1.19

0.005

Ni

1.75

5.160

Pb

0.32

60%) Economic water scarcity (use of water resources < 25%, low-income region) Not estimated

Figure 15.6 Hottest geothermal energy sources around the world which are suitable for desalination, cogeneration and poly-generation schemes [19, 22].

Geothermal water temperature (Quality)

Geothermal Desalination

657

>150 ºC

150 °C) was reported between 2,500 mg/L and 81,000 mg/L while for medium temperature geothermal waters (90–150 °C), a TDS range of 1,100 mg/L 8,200 mg/L. Most of the TDS concentrations were reported between 500 mg/L and 5000 mg/L. A very high concentration range between 260,000 mg/L and 280,000 mg/L was reported for a geothermal plant (Hudson Ranch I and II) in California, U.S.

Geothermal Desalination

15.5.2

659

Geothermal Water for Thermal Desalination

Direct use of geothermal energy is suitable for thermal desalination processes including multi-effect distillation (MED) and multi-stage flash distillation (MSF) process. It is ideal for multi-effect distillation process depending on the source temperature [25]. MED process is preferred due to lower energy requirements in comparison to the MSF process. Generally geothermal energy applications tend to be very site specific and design decisions for one location may not valid for another. MED plants are typically built in units of 2000 to 10000 m3/d (see Figure 15.9). Some of the more recent plants have been built to operate with a top temperature (in the first effect) of about 70°C (158 °F), which reduces the potential for scaling of seawater within the plant but in turn increases the need for additional heat transfer area in the form of tubes. Although the number of MED plants is still relatively small compared to MSF plants, their numbers have been increasing. The cost of an MED plant heavily depends on the performance ratio. Capital and energy costs are significant factors. The main energy requirement is thermal energy. A plant operating with a performance ratio of 8, the thermal energy consumption is around 290 kJ/kg of fresh water and electrical energy demand is 2.5–3 kWh/m3 [25]. The MED process can utilize the low enthalpy geothermal resources (hot geothermal water 80–90 °C) for the production of fresh water. Generally, water costs of less than 1.5 $ per m3 of product water are possible, which make the geothermal-MED coupling very attractive [25]. An MED desalination plan driven by hot geothermal sources with capacity

Feed preheater

Geothermal water

v1

v2

1

Intake seawater vn v3

2

vn-1

3

n-1

Non-condensable gases to-TES Last effect condenser

n Cooling water - TES

Brine reject

Distillate

Condensate D1

D2

Dn-2

Dn-1

Figure 15.9 Multi-effect evaporation desalination system driven by geothermal sources.

660

Desalination 2nd Edition

of 500 m3 of fresh water per day can be considered for understanding the energy requirements. For the evaluation of the desalination plant the following assumption are made: Capacity (distillate produced) D = 500 m3/d Feed seawater TDS concentration W0 = 3.5% Brine TDS concentration WB = 1.25 W0 Brine outlet temperature TB = 40 °C Distillate outlet temperature TD = 40 °C Geothermal water inlet temperature TG,IN = 80 °C Geothermal water outlet temperature TG,OUT = 60 °C Electrical energy consumption 3 kWh/m3 From the mass and energy balance the following results derive: Feed seawater flow rate, M0 = 1500 m3/d Brine flow rate, B = 1000 m3/d Geothermal hot water flow rate, MG = 1875 m3/d Thermal Energy Consumption per kg of distillate: 313.5 kJ/kg

15.5.3

Geothermal Water for Membrane Desalination

Feed water temperature influences the production rates in membrane desalination. Low temperature feed water has higher viscosity and higher resistance to pass through the membrane while high feed water temperature produces high flux (higher production) due to lower viscosity. Membrane processes require higher mechanical energy to pump the cold feed water in winter seasons to meet the daily production rates or higher quantities of feed water needs to be processed. Geothermal waters with high salinity (TDS) can serve as feed water from which freshwater can be produced. An increase in permeate flux of 60% was reported when the feed water temperature was increased from 20 °C to 40 °C [26]. The relation between the permeate flux rate and the feed water temperature is shown in Figure 15.10 [27]. It should be noted that the temperature tolerance of the RO membranes is in the range 20–35 °C. Figure 15.10 inset shows the fraction of the permeate flow that would be affected by the feed water temperature. The permeate flux rate increases by 34% when the feed temperature is increased from 25 °C (1,000 m3/d) to 35 °C (1,344 m3/d) theoretically [26]. In other words, roughly a 6.1% increase in permeate flow rate for every 2 degree temperature difference can be achieved by utilizing process waste heat sources. When process waste heat is not available, utilizing solar collectors or geothermal waters is a feasible option.

Geothermal Desalination

661

2000

Water flowrate (m3/hr)

1800 1600 1400

50

1200

40

1000

30

800 600

10

400

0

200

–0.5

0

0.5

1

0 0

5

10

15

20

25

30

35

40

45

50

Water temperature (ºC)

Figure 15.10 Relationship between the permeate flow and the feed water temperature in a SWRO process. The inset shows the fraction of flow increase with feed water temperature [26].

15.6

Geothermal Desalination – State of the Art

The application of geothermal water in desalination is a relatively unexplored technical concept [28]. Limited number of studies evaluating the potential of geothermal water as a heat source for desalination are available. Here, a brief review on the geothermal energy driven thermal, membrane and hybrid (membrane desalination, humidification and dehumidification) desalination technologies is presented.

15.6.1

Thermal Desalination Processes

The first study of geothermal desalination was proposed and analyzed by Awerbuch et al., in 1976 to produce power and water from geothermal brines in a novel process [38]. In this process, a separator, steam turbine and a MSF unit were included. The separator was used to make sure that the steam flashed from the hot brine extracted from the geothermal production well was circulated in the steam turbine while the non-evaporated hot brine was used as the feed water to the MSF unit to produce fresh water. Another study evaluated the feasibility of geothermal energy driven vertical tube evaporator (VTE), MSF and a high temperature electrodialysis (HTED) desalination process [30]. Smooth and enhanced heat transfer surfaces (tubes) were used in the distillation units while previously tested

662

Desalination 2nd Edition

Teflon-backed membranes, dacron and polypropylene-backed membranes in thin cell configurations were investigated in HTED process. Data was collected to analyze the heat transfer coefficients, tube fouling and scaling effects, feed, brine and product compositions, cell-pair resistance, current efficiency, membrane fouling and scale effects. To control the effects of calcium carbonate, silica, and barium sulfate scaling, a pretreatment consisting of poly-phosphonate addition in distillation units and acidification was done for electrodialysis. MSF and VTE units were tested at high and low temperatures at 270 °F and 190 °F while HTED operation included twostage and three-stage operations at 140 °F and 160 °F respectively. The advantage with geothermal sources is that energy output is generally invariant with less intermittence problems making them ideal for thermal desalination processes. Another possible advantage with geothermal waters is that the feed water itself can be replaced by the geothermal waters; in other words, the geothermal water can serve both as feed and heat transfer medium for desalination. Karystsas has described a case study of a low enthalpy geothermal energy driven seawater desalination plant on the Milos Island in Greece [31]. The proposed design consists of coupling MED units to a geothermal groundwater source with temperatures ranging from 75 °C to 90 °C. The study showed that the exploitation of the low enthalpy geothermal energy would help save the equivalent of 5000 TOE/year for a proposed plant capacity of 600–800 m3/day of fresh water. Even in the case of limited geothermal energy, thermal desalination processes such as MED, thermal vapor compression (TVC), single-stage flash distillation (SF) and MSF can benefit greatly when coupled to geothermal sources by economizing considerable amounts of energy needed for pre-heating. Bouchekima [32] analyzed the performance of solar still using geothermal sources in South Algeria with maximum temperatures of 60–70 °C. A solar distillation (capillary film solar distiller) system was developed and its performance was studied. Theoretical analyses of the heat and mass transfer mechanisms inside this solar distiller were compared with experimental results from the distillation unit. With heat recovery in the capillary film distiller was able to produce up to 20 L/m2/d while a conventional solar still could produce 5–6 L of fresh water per square meter per day of collector surface. Bourouni [33] demonstrated an aero-evapo-condensation process which was found to be promising for cooling as well as for desalting geothermal water. In another study, a geothermal spring with a water temperature of about 70 °C was used to evaluate the heat transfer of air-water-vapor mixtures in the aero-evapo-condensation process [34]. The influence of

Geothermal Desalination

663

different thermal and hydrodynamic parameters on the unit performances was also investigated. The experimental results correlated well with the theoretical predictions. The geothermal desalination costs for this process were compared with other desalination options. The costs for this process were $1.15/m3 when powered by geothermal energy sources; however, when heat is supplied by a fuel source, the cost was $4.80/m3. While the costs for other options were estimated as follows. MSF with back-pressure steam-turbine - $1.57/m3; MSF with gas turbine and waste-heat boiler $1.44/m3; MSF/TVC with gas turbine and waste-heat boiler - $1.31/m3; and reverse osmosis single-stage with energy recovery - $1.39/m3. Mohamed and El Minshawy, conducted theoretical and experimental studies on desalting water using humidification-dehumidification processes similar to Bourouni et al., in Egypt [35]. A humidification and dehumidification was studied by Mahmoudi et al. [36]. In this process, a ground heat exchanger was used to increase the temperature of the geothermal waters acceptable for the proposed desalination process which included evaporative and condensing surfaces while heating the greenhouse. An integrated configuration including a multieffect boiling unit and a MSF unit was evaluated in a feasibility study utilizing geothermal sources in Baja California, Mexico [37]. The geothermal source was available as a heat source at 80°C. In this process, an optimum desalinated water (freshwater) to geothermal source (heat source water) was found to be 1:14. In another study, coastal geothermal desalination plant including MSF-MED configuration was studied. The prototype was developed and tested at a geothermal source temperature of 80 °C. This study reported a much smaller ratio of 1:5.9 for the freshwater to geothermal source. About 20 m3/d of freshwater was produced from a geothermal heat flow of 118 m3/d [38]. Table 15.1 provides a summary of the process conditions for various geothermal installations around the world. This table shows the geothermal water temperatures and the flow rates required in different desalination processes including MED, MSF, vertical tube evaporator, humidificationdehumidification, and membrane distillation. The geothermal water flow requirements and the costs depend on the desalination process configurations.

15.6.2 Membrane Desalination Processes A fluidized bed crystallizer and an air-gap membrane distillation (AGMD) unit were investigated for the suitability of geothermal energy sources by Bouguecha and Dhabi [41]. The membrane surface area in the AGMD was

That study found that the best operating parameters are 85 °C for a feed brine temperature at the evaporator inlet and a circulation flow of about 170 kg/h. Under these conditions, a GOR value of 3.7 and a water production of 16 kg/h may be reached. The integration of one membrane module distiller as a second step at the MED outlet permits an increase of distilled water production by about 7.5% and improvement of the energetic efficiency by practically 10%. T=100 (steam) at 454 kg/h; Freshwater production rate of 18.9 m3/d T=100 (steam) at 3402 kg/h; Freshwater production rate of 79.5 m3/d

MD coupled with multiple effect distiller

MED/VTE* 2 effects

Tunisia

Salton sea/ Imperial valley (USA)

VTE – vertical tube evaporator

MED/VTE* 15 effects

At 75–90 °C geothermal temperatures, 75% energy savings were reported for a geothermal powered desalination process. Produced water cost is estimated as $1.2/m3

Humidification and dehumidification process (HD)

Tunisia

[39, 40]

[41]

[34]

[31]

Geothermal water flow rate of 60 m3 /h at a wellhead temperature of 61–62 °C from a borehole of 188 m deep. Total production of fresh water is 80 m3/day Produced water cost is estimated of the order of 1.6 €/m3 (including only annual operation costs)

References

Kimolos, Greece A two stage MED with distillation under vacuum in vertical tubes

3

[37]

Combination of Multi-Effect Distillation (MED) and Multi-Stage Flash (MSF), called “Multi-Flash with Heaters” (MFWH).

Baja California, Mexico

Process conditions An initial temperature of 150 °C, 4 m of geothermal water was required to produce 1 m3 of desalinated water. At an initial temperature of 80°C, 14 m3 of geothermal water was required.

Process description

Location

Table 15.1 A summary of geothermal desalination installations around the world.

664 Desalination 2nd Edition

Geothermal Desalination

665

64±0.04 cm2 per cell in a three-cell AGMD unit. The membranes were supported by polypropylene grids in 12 mm thick Plexiglas frames. Cooling plates were included to provide condensation of the water vapor. Salt water concentrations were tested between 3 and 35 g/L concentrations and the permeate flux decreased with increase in the salt concentration. This was attributed to vapor reduction due to salt interference effect, increased temperature polarization and concentration polarization. Further, the effect of process parameters such as temperature and feed flow rate were evaluated. An increase of the permeate flow with feed velocity seems to reach maximum value asymptotically to 7.5 kg.m–2h–1. Enhancement of feed water temperature for seawater reverse osmosis plants located in southern California induced a substantial reduction in the cost of potable water [42]. The membrane productivity increase is about 2–3% per 1°C increase of the feeding temperature. Most of the membranes commercialized for RO desalination processes can tolerate temperatures up to 40 °C. However, few membrane suppliers offer new membranes for high-temperature applications. For example, BackpulseableTM membranes (polypropylene tubular membranes) can tolerate temperatures up to 60 °C. It is important to note that an increase of the feeding water of the RO desalination plant of Gabes to 40 °C (temperature tolerance of most commercialized membranes) will increase its productivity of about 20–30%. Boron concentration seems to be an issue with geothermal waters limiting their applications and other beneficial uses. In addition, boron at elevated levels can cause reproductive and developmental toxicity in animals as well as affecting crops. For these geothermal sources, additional boron removal processes must be added to desalination plants. Several studies were reported in Poland and Turkey on various geothermal waters, even containing radioactive compounds such radium, radon and tellurium. In Poland, due to increasing levels of water stress, geothermal waters are increasingly evaluated for their beneficial uses [43]. Several hybrid technologies integrating membranes (microfiltration, ultrafiltration, and reverse osmosis) and ion exchange resins and electrodeinonization processes were evaluated to decrease the Boron concentrations to the acceptable levels (usually below 1.0 mg/L). Desalination of spent geothermal waters for beneficial uses was pursued in Poland. Three different geothermal waters were treated using the double stage membrane process. The sites were Podhale basin (GT-1), Polish Lolands (GT-2) and Western Carpathian Mountains (GT-3). These geothermal waters had a wide range of concentrations of minerals including high iron, strontium, boron and silica [44, 45]. A process schematic of the membrane process is shown in Figure 15.11. The geothermal water characteristics are shown in Table 15.2 [44, 45].

Desalination 2nd Edition

666

NaOH

HCI

RO2

Final treatment

RO1

Geothermal water

Retentate

Pretreatment

Product water

Figure 15.11 Schematic of a double stage RO process for recovering spent geothermal waters for drinking purposes in Poland.

Table 15.2 Geothermal desalination feed and product water compositions studied in Poland and Turkey. Composition

GT-1 Feed

GT-2 Feed

GT-1 Product

GT-2 Product

Na+ (mg/L)

466.8

2297

40.88

151.8

K+ (mg/L)

45.2

27.2

0.83

1.76

196

146.8

CaOH2+

>CO3–

+HN Na+

>CaSO4– +CaOOC

>CaCO3–

Figure 16.11 Electrostatic bonds between calcite surfaces and oil (from Brady and Thyne, 2016).

waterflood. Similarly, equilibrium with barite or gypsum will set upper limits on the amount of sulfate in a waterflood.

16.4.2.2

Unconventionals

In theory, “waterfloods” might be designed for low permeability tight formations that are hydrofractured. Tight formations can’t be waterflooded in the traditional sense because they are essentially impermeable. But they are “waterflooded” when they are fracked. That is, while the goal of the frac fluid injection is to open up fractures and emplace proppant, it also flushes the oil-mineral interfaces in the formation and should affect oil adhesion (e.g. Kurtoglu, 2014), just as occurs during a designer water flood. Working out what frac fluid chemistry (e.g. salinity, hardness) results in the greatest recovery has only recently become a goal. It is driven by: 1. The need to increase the very low recovery from tight formations 5–7%, and; 2. The need to minimize water use. Recall that oil recovery from conventional fields is typically 30% or greater, which is far higher than is recovery from tight formations. Secondary and tertiary recovery (waterflooding and enhanced oil recovery) can boost the amount of oil recoverd from a conventional reservoir to 70% or higher. This is unlikely to happen in unconventionals because of their low permeabilities and the probability

Future Expectations 705 of much of the oil being effectively occluded in the reservoir. But a 30% improvement in unconventional recovery, say from an average of 7% to 10% multiplied over the roughly 4 million barrels of oil produced per day from U.S. unconventionals amounts to nearly $300B per year. Many of the shale plays are in water-limited regions of North America which means there is a strong incentive for achieving the maximum amount of additional oil recovery per gallon of frac water used. We refer to this as “Treat to Need.”

16.4.3

Treat to Need

Much less water is produced from fracked unconventional formations than conventionals, but its salinity can be much higher, sometimes approaching 300,000 mg/L. All of the produced water from conventional reservoirs, and the flowback from tight formations, must either be reused, for example in waterflooding of conventional reservoirs or for subsequent frac jobs (referred to here as “follow-on fracs”), or else it must be deep well injected or evaporated in surface evaporation ponds. Again, excessive deep well injection of produced water occasionally causes seismic damage (e.g. Walsh and Zoback, 2015), so there is motivation to decrease deep well injection whenever possible. The high salinities of frac flowback make them a potential threat to freshwater sources, so there is a desire to treat and reuse the flowback for follow-on fracs. This decreases disposal volumes while freeing up freshwater that would otherwise be used for fracking. But overall, produced water and flowback reuse is done on an ad hoc basis to respond to local freshwater scarcity, or a lack of deep disposal wells. Water reuse decisions should ideally be based as well upon the increase in oil recovery they produce. Indeed, oilfield reuse, rather than treatment for agricultural or non-oil industrial needs, is frequently the most efficient option. High transportation energy costs drive local reuse; the high salinity and NORM in flowback and produced water limit agricultural use. Because the specific chemistry of the water (salinity, Ca+Mg, pH) used in a waterflood or a frac job has a big effect on the amount of oil that is produced, water reuse targets for particular reservoirs need to be matched with the source water and water treatment paths needed to achieve maximal recovery. The reuse recipe will vary from reservoir to reservoir; for uncoventionals, the reuse recipe is also specific to the type of follow-on frac job--slickwater, gel, or hybrid frac. For gel and hybrid fracs, recycled flowback must be treated to the salinity, hardness, etc., that allows the cocktail of gel+surfactants+additives to perform well.

706

Desalination 2nd Edition

16.4.4 Treatment of Hydrofracking Flowback Figure 16.12 shows the locations of the major U.S. shale gas plays and basins. Figure 16.13 shows the basics of the hydrofrack process: 1–3 million gallons of water are mixed with an assortment of HF chemicals (Table  16.4) in stages to: clean the well bore, impose semi-vertical fractures in the shale, and then emplace proppants to hold the fractures open and allow methane to escape. Slickwater is the fluid used to impose the fractures and deliver proppant. Flowback is the water that is pumped back out of the well and is a mixture of shale water (the fluid that was originally in contact with the shale), hydrofrack additives, and residual slickwater. Flowback is held in ponds and then either recycled, injected into Class II UIC wells, or treated and discharged to surface waters. The HF stages include (Additives in parentheses): 1. Acid stage – cleans near wellbore (Acid) 2. Slickwater pad stage – opens the formation (Friction reducer, KCl) 3. Proppant stage – inserts proppant (Proppant, Gelling agent, pH adjusting agent, Crosslinker (early), Breaker (late)) 4. Flushing stage – flushes excess proppant from wellbore.

Figure 16.12 OU.S. Gas plays and basins (from U.S. Environmental Protection Agency 2011).

Future Expectations 707

Figure 16.13 Schematic of hydrofrack process (from U.S. Environmental Protection Agency.

Biocides, Scale inhibitors, and iron control agents are also introduced to prevent the formation of sulfate and carbonate minerals, and iron (hydr) oxide. Fracking is typically done 2,000–6,000 feet below overlying drinking water aquifers which is thought to be sufficiently deep to prevent drinking water contamination. Any future hydrofracking regulations will probably require treatment of flowback salinity to drinking water standards (~ 500 mg/L) if it is to be discharged to surface waters or wastewater treatment plants. The high salinities of the flowback makes discharging it untreated to traditional wastewater treatment plants lethal to their existing biodegradation processes. Deep well injection, when available, is the most effective means for disposing of flowback. Because of its expense, treatment will probably only be pursued where injection wells are absent and, again, to reuse the water for follow on fracs that produce more oil. The key will be to modify traditional treatment steps (e.g. biodegradation, coagulation) to work in the high TDS flowback. Treating flowback water from hydrofracking poses an enormous technical challenge primarily because of the extreme salts loads (200,000 mg/L and higher) that approach halite saturation. Naturally occurring radioactive

708

Desalination 2nd Edition

Table 16.4 Frack fluid additives (after U.S. Department of Energy - All Consulting 2009). Additive

Chemical

Action

Diluted acid

HCl

Helps dissolve minerals and initiate cracks in the rock

Biocide

Glutaraldehyde

Breaker

Ammonium persulfate

Allows delayed breakdown of gel polymer chains

Corrosion inhibitor N,n-dimethyl formamide Crosslinker

Borate salts

Maintains fluid viscosity

Friction reducer

Polyacrylamide, mineral oil

Minimizes friction between fluid and pipe

Gel

Guar gum or Thickens fluid and suspends hydroxyethyl cellulose proppant

Iron control

Citric acid

KCl

Prevents clay swelling

Oxygen scavenger

Ammonium bisulfite

pH adjusting agent

Na2CO3

Proppants

Quartz sand, ceramic

Scale inhibitor

Ethylene glycol

Surfactant

Isopropanol

Corrosion protection Hold fractures open Increases fluid viscosity

material - radium, thorium, uranium, and radon - are also present along with mercury, lead, and arsenic. High sulfate levels complicate treatment further by prompting the formation of sulfate scales containing Ca, Ba, and Ra. Reduced iron is present at the tens of mg/L level in formation waters and can be expected to form ferric hydroxide flocs once exposed to oxygen during treatment. Oil/grease and the frack additives would, from a less saline fluid, be most easily removed by some combination of biodegradation, flocculation, and oxidation (+ filtration). Developing stand-alone treatment steps to break down/remove organics in the face of high TDS is an important research goal. New methods are needed to achieve zero-liquid discharge, the turning of the brine into a sludge (Sludges typically face fewer regulatory hurdles when landfilled than waste brines do when discharged to surface waters). Figure 16.14 emphasizes the fact that traditional water treatment approaches such as EDR, RO, and ion exchange cannot be used on hypersaline flowback. The very high salt loads instead require evaporation and distillation approaches. Figure 16.15 schematically outlines the relative costs

Future Expectations 709 of concentrate management. Keep in mind that the low-cost solutions: surface water and sewer discharge are neither likely to be allowed for flowback disposal. Consider that half a billion gallons of water would be required to dilute a million gallons of 250,000 mg/L flowback down to 500 mg/L.

Bakken frac flowback water Barnett shale produced water

WATER TO BE TREATED

Barnett shale frac flowback water TDS LEVELS ACCEPTABLE FOR TREATMENT

Crystallizer MVR evaporation

BRINE WASTE FROM TREATMENT

Reverse osmosis Electrodialysis reversal

0

0

00 0, 30

25

0,

00

0 0,

00

0 20

15

0,

00

0 10

50

0,

,0

00

00

0

Ion exchange

TDS, ppm

Figure 16.14 Salt levels in Bakken Play fluids (from Energy & Environmental Research Center 2010).

Evaporation pond

LD

/Z or

Capital cost

t

tra

n ce

on

B

ri

tion

c ne

ay

Spr

Deep well injection

a irrig

ch.

ter dis

e wa Surfac

ge

Sewer dischar Concentrate volume

Figure 16.15 Concentrate management costs (After Mickley, 2006).

710

Desalination 2nd Edition

16.4.5 Water Treatment and the Oil Sands The oil sands of Alberta have proven oil reserves of 1.75 trillion barrels, reserves only exceeded by those of Saudia Arabia. The oil is in the form of solid bitumen that is heated with steam to make easier its collection when it is at depth, or mined if it is near the surface. Water is used in both cases – for steaming at depth and for washing sand and clays from surface-mined bitumen. One means for steaming, Steam Assisted Gravity Drainage (SAGD), is depicted in Figure 16.16. Limited water availability has prompted reuse. But steam injected into bitumen-bearing formationsdissolves silica which must be removed before water reuse, otherwise silica (and calcite) scale forms and reduces heat transfer in once-through steam generators limiting their effectiveness and increasing costs. Figure 16.17 shows how warm lime softening is practiced in the SAGD process to remove scale-forming elements. A portion of the boiler blowdown is recycled and combined with silica-rich steam condensate from the free water knockout and low silica makeup water, followed by warm lime softening in a clarifier reactor. The non-recycled portion of the blowdown is typically disposed of by deep well injection. After warm lime softening, filtration, then treatment by weak acid cation exchange resins produces a low hardness, low silica water for further steam generation. Warm lime softening involves adding slaked lime and soda ash to raise the wastewater pH and precipitate hardness as calcite, and addition of magnesium oxide to aid precipitation of silica. A flocculant is used to stabilize and enhance the floc that forms; and a portion of the floc is recirculated to nucleate further precipitation. The process is run at temperatures of 70–90 °C to take advantage of faster reaction rates and the lower solubilities of the metal hydroxide and carbonate minerals that form. A relatively high volume of warm lime softening sludge is produced and ultimately landfilled. Better silica removal approaches would allow greater water reuse and less waste disposal. At the mechanistic level, warm lime softening removes silica as follows: slaked lime (and MgO and soda ash) addition raises the pH; high pH deprotonates silicic acid to silicate ion which combines with a metal cation (Ca+2 or Mg+2 from dissolution of the lime and MgO), or the positively charged surface of MgO, to form a poorly defined amorphous, hydrated metal-silicate mixture. A number of other metals are also known to precipitate/coagulate silica (Wohlberg, Worland et al., 1999). A number of alternatives to warm lime softening exist, such as; 1. Ion exchange to remove all multivalent cations (Portelance 2000; Bridle 2005) would make unnecessary a silica removal

Future Expectations 711 Steam Injector well Oil Recovery well

Sandstone

Steam injection from upper wellbore Heated bitumen flows to lower wellbore

Figure 16.16 Schematic of SAGD process (source: http://pubs.usgs.gov/fs/fs070-03/fig3. jpg).

SAGD process

Makeup

Reservoir

Oil removal

Inputs CaO + MgO + coagulant

Outputs Warm lime softening

Filtration

WAC Steam to reservoir

Once through steam generator

Sludge

Backwash

Regenerator waste

Backwash

Figure 16.17 Warm lime softening treatment for SAGD.

Desalination 2nd Edition

712

step because silica scale formation requires multivalent cations to accelerate polymerization, coagulate silica polymers, and form metal silicates. Ion exchange should therefore allow much higher silica waters to be used in steam generators. 2. Mechanical Vapor Compression (MVC) produces an essentially pure feedwater and a blowdown that is ~ 33-fold more concentrated than the makeup water (e.g. Portelance 2000; Heins 2010) and which must be treated before deepwell injection, or evaporated using a crystallizer to a cake for zero liquid discharge landfilling (e.g. Heins, McNeill et al., 2005). Silica removal is a central requirement of MVC concentrate management (e.g. Portelance 2000). 3. Silica scale inhibitors might allow higher silica levels to be present during steam generation (Pedenaud, Goulay et al., 2006). Lowering solution pH might also apply. The primary drawbacks to warm lime softening are extensive chemical inputs (lime, soda ash, MgO, coagulants, regenerants for the ion exchange step that follows warm lime softening) and management of the high water sludges that are produced. An important advantage of warm lime softening is the long industrial experience with the process. Also, managing warm lime softening sludges aboveground might be more easily done than assuring the sustainable injection of e.g. MVC concentrate at depth. Oil sands that are processed ex situ pose an even larger water treatment challenge. Oil sands are washed in a dilute base to loosen the bitumen from the sand. The resulting wastewater includes toxic naphthenic acids and mineral fines – clay particles that settle very slowly and prevent the rapid reuse and/or disposal of the water. The wastewaters are piped to areally extensive settling ponds that pose a health threat to wildlife. Settling/coagulating the clay fines rapidly would allow much smaller ponds. Presently, a slurry of gypsum provides some settling of fines. Better methods are needed though. Note that any treatment alternative must be relatively inexpensive because of the very large water volumes involved.

16.5

The Future of Desalination Technologies

Seawater RO presently requires an energy expenditure in the range of 9–60 kJ/kg at a cost of $2 to $4 per 1000 gallons of water produced. The

Future Expectations 713 Fixed charges-37% Membrane replacement-5% Labor-4% Maintenance & parts-7% Consumables-3% Electric power-44%

Figure 16.18 Cost breakdown for seawater reverse osmosis (from Miller 2003).

theoretical minimum energy expenditure is 3–7 kJ/kg (e.g. Miller 2003). The fact that energy consumption is nearly half the overall water production costs (see Figure 16.18) means that order or magnitude reductions in desalination costs probably won’t occur in the future but that low energy desalination will be an important goal for the future. The need forimprovements in multiple areas is emphasized in the research targets identified in the National Desalination Roadmap Implementation Report (American Water Works Association Research Foundation, WateReuse Foundation et al. 2010). The research targets can be grouped into Membranes, Concentrate Management, and Alternative Technologies. Membranes 1. Seawater pretreatment conditioning optimization, 2. Integrated treatment approaches for desalination-specific problem contaminants (e.g. boron, NDMA). 3. Alternatives to spiral membranes/spiral membrane configuration optimization, 4. Mechanistic understanding of high-pressure membranes, 5. Pretreatment to prevent membrane biofouling, 6. Boron removal by membranes, 7. Improved feedwater recovery for membrane desalination, 8. Fouling characterization, sensing, and prediction, 9. Post-treatment stabilization of desalinated water, 10. Variable operating rate capability desalination plants, 11. Biomimetic desalination, Alternative Technologies 1. Hybrid reverse osmosis/electrodialysis desalination, 2. Membrane distillation,

714

Desalination 2nd Edition 3. Forward osmosis, 4. Nanostructured high flux membranes,

Concentrate Management 1. 2. 3. 4. 5. 6.

Hybrid techniques for zero liquid discharge, Self-sealing evaporation ponds, Evaporation enhancement (for concentrate management), Selective contaminant removal from concentrate for resale, Silica removal, Selective precipitation of concentrate salts,

Other 1. Minimization of anti-scalants and biocides, 2. Energy recovery from desalination, and 3. Co-siting of desalination with wastewater treatment, power generation, oil and gas, and other industrial facilities. A number of the above have been touched upon in this chapter, including: silica removal, anti-scalants and biocides, selective precipitation of concentrate salts, and co-siting of desalination facilities. Energy recovery is discussed in the chapter by Pankratz. Forward osmosis is examined in the chapter by McCutcheon. Membrane distillation is covered in the chapter by Cath. Electrodialysis and electrodeionization are explored in the chapters by Moon and Lee and Wood and Gifford. Many of the research targets involve biofouling are discussed in the several chapters. And the remainder of non-traditional desalination approaches (e.g. clathrates, wind and wave-driven desalination, osmotic pumps, offshore desal, fog collection, solar stills, centrifugal desalination) have been described in detail by Miller (2003). Improvements in RO technology focus on high flux, low energy membranes in particular biomimetic desalination and nanostructured membranes, and supporting efforts in computational materials chemistry. We anticipate that national and international scientific initiatives in nanotechnology and biotechnology will result in improved desalination technologies in the future. Nanotechnology—the use of materials and processes that operate over a length scale of roughly 1 to 100 molecular diameters— is the target of over $10 billion in R&D investment each year. Global biotech R&D investments amounted to roughly $30 billion in 2010. While a relatively small fraction of these efforts focus on water treatment, the broad scale technical advances that are occurring in both fields are expected to benefit desalination(e.g. Shannon, Bohn et al. 2008).

Future Expectations 715 A key goal of desalination research is to build a molecular level understanding of water behavior at membrane interfaces and use it to precisely design high-flux membranes (Cygan, Brinker et al. 2008). Membranes might be designed that prevent attachment of biofilms, or that don’t degrade in biofilm-destroying oxidants. The expectation is that parallel advances in materials development will make the transition from the lab to the market rapid for each technology. Note thought that existing processes, such as coagulation, lime softening, micropollutant filtration, and ion exchange for example, already rely on nanoscale processes. By the same token the design of existing RO membranes involves an ample contribution of nanoscale coordination chemistry. Recent emphasis on better theoretical understanding and resolution of these processes at the molecular level is what is expected to drive the respective technologies to higher levels of performance and lower costs.

16.5.1 Biomimetic and Nanotech Membranes Examples of biological desalination include mangroves, which separate H2O from seawater and secrete salt crystals in their roots, and seabirds far from shore that separate water in a gland above their beak and sneeze out the concentrate. Peter Agre received the Nobel Prize in Chemistry in 2003 for his co-discovery of aquaporins “the plumbing system for cells”. Cell walls of most organisms contain specialized protein channels, ‘aquaporins’, that quickly transport water or ions selectively across the cell membrane (Borgnia, Nielsen et al. 1999). Water is transported in a hydrophobic channel a few Ångstroms in diameter, while hydrogen bonding interactions with functional groups on the channel walls preferentially orient the molecules in single-file fashion inside the pore. Ions are effectively excluded because of the large increase in free energy associated with penetrating a channel due to loss of its hydration sphere and the associated decrease in entropy. The effectiveness of aquaporins in shuttling water through cell membranes has motivated the search for aquaporin-assisted membranes, and for synthetic analogues. While natural aquaporin proteins extracted from living organisms can be incorporated into a lipid bilayer membrane or a synthetic polymer matrix, porous inorganic membranes modified to provide aquaporin-like function may provide a more robust alternative. These include carbon nanotubes, double-walled carbon nanotubes, and metal oxide frameworks. Making carbon nanotubes is presently an involved technical process. Typically a substrate containing metal seeds of the same diameter as the nanotubes are heated to 600 to 900 oC, and then

716

Desalination 2nd Edition

a carbon-containing gas such as methane or alcohol is added. Nanotubes then grow from the metal seeds. The metal from the seeds are problematic in that the metal can later occlude nanotubes. Large-scale manufacture of nanotube-based membranes will require vastly improved manufacturing processes to be economical.Self-assembly and template directed synthesis of aquaporin-simulating materials is one potential way out.

16.5.2 Desalination with Renewables Desalination is a particularly energy-intensive form of water treatment. Less than 1% of water that is desalinated uses renewable energy, in part because renewable desalination is still more expensive than desalination powered by fossil fuels (International Renewable Energy Association, 2012). Renewable desalination is cost competitive in remote regions though. Solar photovoltaics (PV) are used to drive RO-based desalination. The world’s largest solar PV desalination plant is being built in Al Khafji in Saudi Arabia, which will rely on membrane filtration to produce 60,000 m3 of water per day. Concentrated solar power is being considered for solar thermal desalination. Wind-powered RO units have been installed at several coastal locations. Geothermal desalination remains less used. Because renewable energy output and water usage vary across the course of a day, there is the potential for “storing” renewable energy as treated water.

16.6 Summary As presented in this chapter, supplies of fresh water have become limited in many regions of the work and the U.S. due to economic growth and development. To meet the growing water needs of industrialized and developing nations, increased efficiency in fresh water use and increased utilization of impaired waters through advanced water treatment processes will be required. While the use of impaired or nontraditional waters, such as waste water reuse and desalination started to grow significantly in the U.S. in the 1990s, that trend started in many other regions of the world in the 1970s. All indications are that this trend in the treatment and use of impaired waters to supplement fresh water resources and to better utilize non fresh water resources where applicable and fresh water is not needed, such as in many industrial applications, will continue to grow as demands on the worlds fresh water resources continue to grow. Therefore, development of approaches to more efficiently and effectively treat and utilize impaired

Future Expectations 717 waters such as seawater, brackish ground water, oil and gas produced water, and domestic and industrial waste water, will continue to grow and expand. Many of the impaired and nontraditional waters mentioned have increased levels to even very high levels of salts that will need to be treated, meaning that desalination technologies and improvements in desalination approaches and processes will be vital to enable the U.S. and most other countries in meeting future global sustainable water resource and supply needs.

References 1. Altman, S. J., R. P. Jensen, et al. (2011). “Nanofiltration Treatment of SideStream Cooling Tower Water for Reduction of Water Usage.” Desalination 285, 177–183, 2012. 2. American Water Works Association Research Foundation, WateReuse Foundation, et al. National Desalination Roadmap Implementation Report http://www.sandia.gov/water/docs/DesalImplementRoadmap1–26-2010_c_ web.pdf, 2010. 3. Austad, T., A. RezaeiDoust, et al. “Chemical mechanism of low salinity water flooding in sandstone (SPE 129767).” SPE Improved Oil Recovery Symposium: 1–17, 2010. 4. Borgnia, M., S. Nielsen, et al. “Cellular and molecular biology of the aquaporin water channels.” Annu. Rev. Biochem. 68, 425–458, 1999. 5. Bourcier, W. L., M. Lin, et al. Recovery of minerals and metals from geothermal fluids https://e-reports-ext.llnl.gov/pdf/324646.pdf. Livermore, CA, Lawrence Livermore National Laboratory, 2005. 6. Brady, P. V. and J. L. Krumhansl. “A Surface Complexation Model of Oil-BrineSandstone Interfaces at 100 oC: Low Salinity Waterflooding (81. 171–176.)” Journal of Petroleum Science and Engineering, 2012. 7. Brady, P.V. and Krumhansl, J.L. pH adjustment of power plant cooling water with flue gas/fly ash. US Patent 9140145, 2015. 8. Brady, Patrick V., and Geoffrey Thyne. “Functional wettability in carbonate reservoirs.” Energy & Fuels 30, no. 11, 9217–9225, 2016. 9. Brady, Patrick V., Charles R. Bryan, Geoffrey Thyne, and Huina Li. “Altering wettability to recover more oil from tight formations.”  Journal of Unconventional Oil and Gas Resources 15, 79–83, 2016. 10. Buckley, J. S., Y. Liu, et al. “Mechanisms of wetting alteration by crude oil.” SPE Journal, March 1998, 54–61, 1998. 11. Buckley, J. S., K. Takamura, et al. “Influence of electrical surface charges on the wetting properties of crude oil.” SPE Reservoir Engineering 4, 332–340, 1989. 12. Collins, I. R., G. R. Jerauld, et al. INTERNATIONAL PATENT WO 2008/029131 A1/Hydrocarbon Recovery, 2008.

718

Desalination 2nd Edition

13. Costa, L. M. and P. J. McCabe, US Patent 7514001 B2 High recovery reverse osmosis process and apparatus, 2009. 14. Cygan, R. T., C. J. Brinker, et al. “A molecular basis for advanced materials in water treatment.” MRS Bulletin, 33, 42–47, 2008. 15. Dubey, S. T. and P. H. Doe. “Base number and wetting properties of crude oils.” SPE Reservoir Engineering, August: 195–199, 1993. 16. Energy & Environmental Research Center. Bakken Water Opportunities Assessment - Phase 1; Summary Report http://www.nd.gov/ndic/ogrp/ info/g-018–036-fi.pdf. Grand Forks, ND, Energy & Environmental Research Center, 2010. 17. EPRI . Use of Degraded Water Sources as Cooling Water in Power Plants, California Energy Commission, 2003. 18. Erzuah, S., Fjelde, I. and Omekeh, A.V., 2017, June. Wettability Estimation by Surface Complexation Simulations. In SPE Europec featured at 79th EAGE Conference and Exhibition. Society of Petroleum Engineers. 19. Featherstone, J. L. Geothermal plant silica control system. US Patent 4429535, 1984. 20. Government Accountability Office. Fresh Water Supply: States’ Views of How Federal Agencies Could Help Them Meet the Challenges of Expected Shortages, USGA, ed., Report # GAO-03–514. Washington, DC, p. 110, Jul/2003, 2003. 21. Gritzuk, M. Planned desalination programs by the City of Phoenix. Bi-National Desalination Conference 2004, 2004. 22. Hutchinson, W. Desalination of brackish groundwater and deep well injection of concentrate in El Paso, Texas http://www.gwpc.org/meetings/uic/2009/proceedings/Hutchison,%20Bill.pdf. Groundwater Protection Council, 2009. 23. International Renewable Energy Association, Water desalination using renewable energy: Technology Brief I12, 2012. 24. Kurtoglu, B. Integrated Reservoir Characterization and Modeling in Support of Enhanced Oil Recovery from Bakken, Petroleum Engineering. Colorado School of Mines, Golden, CO, 2014. 25. Lager, A., K. J. Webb, et al. Low salinity oil recovery - an experimental investigation; SCA2006–36, 2006. 26. Lott, R. Water and energy in Arizona http://www.gwpc.org/meetings/ forum/2011/Proceedings/pdf_presentations/9f_Lott_Bob.pdf. Groundwater Protection Council, 2011. 27. Mattson, J. V. and T. G. I. Harris. “Zero Discharge of Cooling Water by Sidestream Softening.” Water Pollution Control Federation, 51, 2602–2614, 1979. 28. Mickley, M. E. State of the Science Report to the Joint Water Reuse and Desalination Task Force. Concentrate Management, 2006. 29. Mickley, M. US municipal desalination plants: number, types, locations, sizes, and concentrate management practices. IDA Journal of Desalination and Water Reuse, 4, 44–51, 2012.

Future Expectations 719 30. Miller, J. E. Review of Water Resources and Desalination Technologies http:// www.sandia.gov/water/docs/MillerSAND2003_0800.pdf. Albuquerque, NM, Sandia National Laboratories, 2003. 31. MIT. The Future of Natural Gas. http://web.mit.edu/mitei/research/studies/ documents/natural-gas-2011/NaturalGas_Report.pdf, 2011. 32. Poulson, T. Future Trends Analysis (Draft Phase I Document). U.S. Department of Interior, Bureau of Reclamation, Central Arizona Salinity Study, 2003. 33. Qiao, C., Johns, R. and Li, L., Modeling low-salinity waterflooding in chalk and limestone reservoirs. Energy & Fuels, 30(2), pp. 884–895, 2016. 34. Rezaei Doust, A., T. Puntervold, et al. “Smart water as wettability modifier in carbonate and sandstone: A discussion of similarities/differences in the chemical mechanisms.” Energy & Fuels, 23, 4479–4485, 2009. 35. Sandia National Laboratories. Energy-Water Challenges and Research and Development Issues. Albuquerque, NM, Sandia National Laboratories, 2009. 36. Shannon, M. A., P. W. Bohn, et al. “Science and technology for water purification in the coming decades.” Nature, 452, 301–310, 2008. 37. Stanton, Jennifer S., David W. Anning, Craig J. Brown, Richard B. Moore, Virginia L. McGuire, Sharon L. Qi, Alta C. Harris et al. Brackish groundwater in the United States. No. 1833. US Geological Survey, 2017. 38. Stepan, D. J., R. E. Shockley, et al. Bakken water opportunities assessment Phase 1; 2010-EERC-04–03; http://www.nd.gov/ndic/ogrp/info/g-018–036-fi. pdf, 2010. 39. U.S. Department of Energy - All Consulting. Modern Shale Gas Development in the United States: A Primer, 2009. 40. U.S. Environmental Protection Agency. Draft plan to study the potential impacts of hydraulic fracturing on drinking water resources, 2011. 41. Walsh, F.R. and Zoback, M.D. Oklahoma’s recent earthquakes and saltwater disposal. Science advances 1, e1500195, 2015.

List of Contributors Chapter 1 Jane Kucera (volume editor) Jane Kucera is a chemical engineer with over 37 years’ experience in the area of membrane technology for water treatment. Jane began her work with membranes in 1982 while working as a Graduate Student in the Seawater Laboratory (today known as the WaTeR Center) at the University of California, Los Angeles, USA (UCLA), where she received her M.S. Degree in Chemical Engineering in 1984. She then went to work for Bend Research, Inc. (BRI), as a research engineer, where she spent 7 years developing novel membrane materials and process designs, including serving as lead researcher for several NASA research contracts to develop membranebased, water recovery and reuse systems for the International Space Station. Later, Jane joined USFilter Corporation, where she was the process design leader for the 2 MGD Honouliuli Water Reclamation Facility for Oahu, Hawaii; the system treats secondary municipal effluent with microfiltration and reverse osmosis desalination for reuse as boiler feed for refineries and power generation facilities on the Island. Jane is currently a Global Senior Technical Consultant in the Engineering and Managed Operations Group at NALCO Water/an Ecolab Company, where she is responsible for developing industrial desalination system process designs. Jane has authored the book, Reverse Osmosis: Design, Applications and Processes for Engineers, in its second edition (2015), and has contributed chapters to several other volumes covering membranes, reverse osmosis, and desalination. She is a member of American Institute of Chemical Engineers (AIChE); Engineers’ Society of Western Pennsylvania (ESWP), where she was appointed to the Executive Committee for the International Water Conference (IWC); American Society of Mechanical Engineers (ASME), where she was appointed to the Research Committee on Water Management Technology and elected vice-chair of the committee; North American Membrane Society (NAMS); International Desalination

Jane Kucera (ed.) Desalination 2nd Edition, (721–735) © 2019 Scrivener Publishing LLC

721

722 List of Contributors Association (IDA); and the Society of Women Engineers (SWE). Jane also received a BA in Chemistry, Summa Cum Laude, from Linfield College in McMinnville, Oregon, USA. Chapter 2 Joachim Gebel Dr. Joachim Gebel received his Doctorate in Mechanical Engineering, Summa Cum Laude, from RWTH Aachen University in 1990. For more than 20 years, Joachim has worked as a consultant in different functions; since 2003, he is co-owner and acting partner of S.T.E.P. Consulting GmbH, Aachen, Germany. Joachim has focused on topics such as seawater and brackish water desalination, waste water treatment, dual-purpose plants and cogeneration throughout his career. With their book “An Engineer’s Guide to Seawater Desalination,” published in 2008, Joachim and his coauthor Dr. Süleyman Yüce reported on important topics of main desalination technologies, from the basics up to material problems, as well as the design and the operation of facilities. Joachim is currently a Professor in the Department of Technology and Bionics at the Hochschule RheinWaal in Kleve, Germany, where he presents lectures on thermodynamics, process engineering and plant design. Chapter 3 Mark Wilf Dr Mark Wilf has been involved in process development, system design, project execution, plant operation and maintenance of large, commercial desalination plants in US, Europe and Middle East since 1977. Mark is a regular contributor to professional journals, wrote chapters on membrane technology processes and applications to a number of books. He edited and wrote with other coauthors The Guidebook to Membrane Desalination Technology that has been published in 2006. The second book on membrane technology he edited and contributed to: The Guidebook to Membrane Technology for Wastewater Reclamation has been published in 2010. Mark regularly presents and teaches desalination and membrane technology subjects to engineers and water professionals. His teaching activity includes teaching course on membrane technology and desalination for the European Desalination Society at the Genoa University, Italy and in San Diego. Chapter 4 Christopher Bellona Dr. Christopher Bellona is an Assistant Professor in the Department of Civil and Environmental Engineering at Clarkson University. Dr. Bellona’s

List of Contributors

723

past research has largely been focused on membrane science and the potable reuse of wastewater effluent. His research has spanned both fundamental and applied aspects of water reuse including the fate and removal of constituents of concern (e.g., organic contaminants, nutrients), membrane fouling, modeling, and pilot-scale testing of membrane technologies. His current research is focused on membrane systems for anaerobic wastewater treatment, extraction of energy and valuable materials from waste streams, and development of novel and hybrid membrane systems. He received his M.S. and Ph.D. degrees in Environmental Science and Engineering from the Colorado School of Mines, and his B.S. in Environmental Science from Western Washington University. Chapter 5 Jeff McCutcheon Jeffrey McCutcheon is an Associate Professor at the University of Connecticut in the Department of Chemical & Biomolecular Engineering. He is also the Executive Director of the Fraunhofer Center for Energy Innovation at the University of Connecticut. Dr. McCutcheon’s research is in membrane separations with a focus on membrane fabrication and characterization. His group focuses on applications in liquid separations, including forward osmosis, membrane distillation, nanofiltration, and organic solvent separations. He has written 65 refereed publications, 3 patents, and 2 book chapters. He has won numerous awards including the FRI/John G. Kunesh Award from the AIChE Separations Division (2014), The DuPont Young Profess Award (2013), the 3M Nontenured Faculty Award (2011), and the Solvay Specialty Polymers Young Faculty Award (2011). As of 2018, he is President-elect of the North American Membrane Society and recently finished his term as Area Chair of Area 2D of the AIChE Separations Division (2015-2017). Lingling Xia Lingling Xia is a Ph.D. student at the University of Connecticut in the Department of Chemical & Biomolecular Engineering, under the advisement of Dr. Jeffrey McCutcheon. Her research is about membrane separation with a focus on membrane fabrication and characterization. Her current work is interested in nanofiltration membrane design for aqueous and non-aqueous solvent system, which has wide range of applications such as wastewater treatment, water softening, and product separation/ purification in pharmaceutical and petrochemical industry. Previously, her work involved the process study of pressure retarded osmosis (PRO), and draw solution design for PRO, specifically for osmotic heat engine.

724 List of Contributors Before landing in UConn, she earned her bachelor degree in Chemical Engineering at Huazhong University of Science and Technology, where she did research about pervaporation membrane design and characterization. Ngoc Bui Ngoc Bui is currently a postdoc researcher at the Molecular Foundry at Lawrence Berkeley National Laboratory (LBNL). Her research interests lie at the interface of nanomaterials and membrane science and engineering for sustainable water, energy and life sciences. Prior to LBNL, she was a postdoc fellow at the Physical and Life Sciences Directorate at Lawrence Livermore National Laboratory (LLNL). At LLNL, she joined a team striving to develop smart dynamic fabric materials that can sense and protectively respond to chemical and biological agents for protective clothing. Dr. Bui earned her PhD in Chemical Engineering at the University of Connecticut in December 2013 (under the advisement of Dr. Jeffrey McCutcheon). Her PhD research focused on pioneering the development of a new and effective generation of nanofiber-supported thin film composite membrane platform for sustainable water and energy production by harnessing salinity-gradient energy via osmotically-driven membrane processes. Her work also focused on developing as thorough a mathematical transport model as possible to deconvolute all mass transfer resistances in water and solute transport across the membrane, offering opportunities for better predictive models and advanced membrane and system designs with precise control to be proposed. Chapter 6 Seung-Hyeon Moon Seung-Hyeon Moon received his B.S. and M.S. in Chemical Engineering from Seoul National University, and Ph.D. in Chemical Engineering from Illinois Institute of Technology, Chicago. Subsequently he joined Energy Systems Division, Argonne National Laboratory in U.S to develop fermentation and bio-separation processes using electromembranes. Since then, he has been working on synthesis, characterization, and process application of electromembranes. In 1994, he moved to Gwangju Institute of Science and Technology (GIST), Gwangju, Korea. He continued the membrane studies for cleaner technologies and industrial wastewater treatment. Further his research areas have been expanded to application of ion exchange membranes for energy conversion processes including fuel cells, dye-sensitized solar cells, rechargeable battery and aqueous and non-aqueous redox flow batteries based on fundamental electrochemistry

List of Contributors

725

and nanoionics to innovate the relevant technologies, including potential measurement in a boundary layer, through-plane conductivity device, and ion channel formation. He is teaching electrochemistry, water chemistry, diffusion, physicochemical hydrodynamics for undergraduate and graduate students. He published over 200 papers and advised over 40 graduate students. He is a fellow of Korea Academy of Science and Technology. In administration of the Institute, he was the founding chairman of School of Environmental Science and Engineering, GIST (1994-1998), and served as Programme Cordinator of GIST-United Nations University Joint Programme on Science and Technology for Sustainability (2000-2004), Dean of Academic and Student Affairs (2004-2006), Vice-President (20062008), Acting President (2007-2008). Also he managed the Energy and Environmental Technology Program in National Research Foundation of Korea (2014-2015). Since Feb. 2015, he serves as the President of GIST. Jae-Hwan Choi Jae-Hwan Choi received his B.S. in Environmental Engineering from Inha University, and M.S. and Ph.D. in Environmental Science and Engineering from Gwangju Institute of Science and Technology (GIST), Gwangju, Korea. He studied the separation technology of ionic materials using ion exchange membranes in his doctoral thesis. The transport phenomenon at the interface of the ion exchange membrane was studied according to the characteristics of the ion exchange membrane and current density. After his doctorate, he moved to the Department of Chemical Engineering at Kongju National University, Cheonan, Korea. Since then, he has been working on the capacitive deionization (CDI) that uses porous carbon electrodes to removes ions by electrostatic attraction. His research group has developed a highly efficient carbon electrode that improves adsorption capacity and charge efficiency by adding functional materials. They have also developed composite carbon electrodes that can selectively remove specific ions from mixed solutions. The developed carbon electrodes are commercially produced by a company and applied to actual desalination processes. Currently, he is continuing his research to develop stable carbon electrodes against electrode reactions. Through these studies, he is expected to contribute to the commercialization of CDI technology, which is emerging as a promising desalination technology. He is teaching electrochemistry, analytical chemistry, environmental engineering, and chemical engineering experiments for undergraduate students. He has published more than 60 research papers on ion exchange membrane and CDI related research.

726 List of Contributors Chapter 7 Jonathan Wood Jonathan Wood is Global Lead, Applications & Technical Support (Electrodeionization and Electrodialysis) for Evoqua Water Technologies in Lowell, MA. He has a B.S. in Chemical Engineering from Worcester Polytechnic Institute in Worcester, MA and an M.S. in Environmental Engineering from Northeastern University in Boston, MA. Jon has over 40 years of experience in water purification, encompassing research, product development, product management and technical support. He started his career with Barnstead Company, working with distillation, pure steam generation, and deionization equipment, and then joined Millipore Corporation, where he was part of the team that commercialized electrodeionization. He has continued to work in this field with USFilter, Veolia, Siemens and Evoqua. Joseph Gifford Joseph Gifford recently became Vice President of Research & Development for Nanostone Water in Waltham, MA, focused on developing ceramic membrane technology. Prior to that he was VP of R&D for Evoqua Water Technologies in Lowell, MA. He was with Evoqua for over 23 years where he held roles in R&D applications and process development, heavily focused on electrochemical technologies such as electrodeionization and electrodialysis, and has a number of U.S. and international patents for technologies he’s helped bring to market. He has a B.S. in Chemical Engineering from Worcester Polytechnic Institute in Worcester, MA and an M.S. in Chemical Engineering from the University of Massachusetts in Lowell, MA. Chapter 8 Xing Yang Dr Yang has been a lecturer at the College of Engineering & Science and a researcher at the Institute for Sustainability & Innovation, Victoria University, Melbourne Australia since 2013. Prior to joining Victoria University, she obtained her PhD in late 2012 from from Civil & Environmental engineering at Nanyang Technological University (NTU), Singapore, then worked as a research fellow at the Singapore Membrane Technology Centre (SMTC) from 2012 to 2013. Dr Yang has over 10 years experiences on membrane research and chemical process design over 10 years in China, Southeast Asia and Australia, and published over 50 peer-reviewed articles in top journals and conferences

List of Contributors

727

in her field. She has been dedicated to translational research and making an impact in the commercialization of membrane processes in industry, including her significant contribution to the development of resource recovery from waste waters using membrane distillation and other hybrid membranes processes. Dr Yang’s current research interests include 1) Membrane process design for wastewater treatment, desalination, zero liquid discharge and resource recovery; 2) Functional material and interfacial modification for anti-fouling properties and functional separation; 3) Transport models and process simulations, cost and energy evaluation of membrane systems. Dr Yang’s research has been recognized internationally indicated by the awards and honours that she received from international conferences and organizations such as the Water Ambassador of Public Utilities Board of Singapore, Young Academics Award by European Membrane Society and the Vice Chancellor Award for research and research training at Victoria University. Dr Yang also actively contributes to the professional and scientific committees, such as: Editorial board member of Desalination Journal; Editor of Heliyon Open Access Journal published by Elsevier Peer reviewer of several top journals such as Journal of Membrane Science. Asia-Pacific regional coordinator of the International Desalination Association (IDA) Young Leaders Program between 2013 and 2017, for facilitating the creation of international young water professional network in water and membrane community; Event coordinator in the committee of the Early Career and Student Symposium of the Membrane Society of Australasia (MSA) 2014; Scientific committee of the organizing team of the International Conference of Aseanian membrane Society (AMS11) 2018; Committee member of the International  Conference  on Desalination using Membrane Technology (MEMDES4) in 2019; Anthony G. Fane Tony Fane is a Chemical Engineer from Imperial College, London, who has worked on membranes since 1973. His current interests are in membranes

728 List of Contributors applied to environmental applications and the water cycle, with a focus on sustainability aspects of membrane technology, including desalination and reuse. He is a former Director of the UNESCO Centre for Membrane Science and Technology at UNSW. Since 2002 he has directed membrane research in Singapore as Temasek Professor (2002-2006) at Nanyang Technological University, then as Foundation Director of the Singapore Membrane Technology Centre (SMTC) at NTU. He is now DirectorMentor at SMTC. The SMTC has a group of over 85 researchers dedicated to fundamental and applied research into membranes for the water cycle. He is on the Advisory Board of the Journal of Membrane Science (former editor from 1992 to 2005) and Editorial Board of Desalination. He is the Patron of the Membrane Society of Australasia. Rong Wang Dr. Rong Wang is currently an Associate Professor in the School of Civil and Environmental Engineering, Nanyang Technological University, Singapore. She also holds the position of Director, Singapore Membrane Technology Centre (SMTC). Dr. Wang has 25 years of experience in Chemical Engineering, Water, Environment and Energy related R&D. She specializes in novel membrane development for water and wastewater treatment, liquid purification and gas separation, and in development of novel hybrid membrane systems and process simulation. Her research interests cover (1) novel forward osmosis (FO)/pressure retarded osmosis (PRO) hollow fiber membranes, low pressure nanofiltration (NF) hollow fiber membranes, bio-mimetic membranes, hydrophobic hollow membranes and mixed matrix membranes for seawater desalination, wastewater treatment, membrane contactor for CO2 capture, etc; (2) simulation and optimization of various membrane processes such as membrane contactor, membrane distillation and separation of mixed gases in membrane modules, etc. Dr. Wang has over 130 SCI-tracked journal publications (cited times: >2600, h-index: 32 @Scopus) and over 100 conference contributions. She is the inventor of 20 patents/ technical disclosures for novel membrane fabrication. She has co-authored 6 book chapters. She is the editorial board member for two journals of Journal of Membrane Science, and Desalination. She was a Guest Editor for two special issues of Desalination in 2011 and 2013. As PI and Co-PI of projects, she has been successfully awarded research grants of over S$15 millions by the governmental agencies and industry in the past five years.

List of Contributors

729

Chapter 9 John H. Lienhard V John H. Lienhard V is the Abdul Latif Jameel Professor of Water and Mechanical Engineering at MIT. During three decades on the MIT faculty, Lienhard’s research and educational efforts have focused on  heat and fluid flow, water purification and desalination, and thermodynamics. Lienhard received his bachelor’s and master’s degrees in thermal engineering from University of California, Los Angeles in the Chemical, Nuclear, and Thermal Engineering Department. He joined MIT in 1988 immediately after completing his PhD in the Applied Mechanics and Engineering Science Department at the University of California, San Diego. Lienhard’s research on water purification has included humidification-dehumidification desalination, membrane distillation desalination, forward and reverse osmosis, batch process design, solar-driven desalination, bubble columns, fouling and scaling, nanofiltration, electrodialysis, management of high salinity brines, thermodynamic and energy efficiency analysis of desalination cycles. Lienhard has directly supervised more than 85 graduate theses and postdoctoral associates and has authored more than 200 peerreviewed publications and several textbooks. He is the founding director of the Abdul Latif Jameel Water and Food Systems Lab, founding director of the Center for Clean Water and Clean Energy, and has directed the Rohsenow Kendall Heat Transfer Lab since 1997. Lienhard is a recipient of the 1988 National Science Foundation Presidential Young Investigator Award, the 1992 SAE Teetor Award, a 1997 R&D 100 Award, the 2012 ASME Globalization Medal, the 2015 ASME Heat Transfer Memorial Award, and several teaching awards. He is a Fellow of ASME and a Registered Professional Engineer. Lienhard and his students received thirteen best paper and poster awards for their desalination research during 2011-2017. He is a cofounder of Gradiant Corporation, which focuses on difficult areas of industrial water treatment. In addition, Lienhard holds more than forty US Patents and pending US applications, many of which have been commercialized in the water industry. Chapter 10 M. Shafiur Rahman Professor Shafiur Rahman is the author/co-author of over 300 technical articles including more than 140 journal papers and 12 books. He is the author of the internationally acclaimed and award-wining Food Properties Handbook, and editor of the popular Handbook of Food Preservation published by CRC Press, Florida. First editions received the bestseller

730 List of Contributors recognition and the third edition will be released soon. First edition was also translated into Spanish. He has initiated the International Journal of Food Properties and serving as the founding Editor-in-Chief for more than 20 years, and serving in the editorial boards of 10 international journals and book series. In 1998 he has been invited and continued to serve as a Food Science Adviser for the International Foundation for Science, Sweden. He was invited as a keynote/plenary speaker for more than 10 international conferences in the food science and engineering area. In 2014, he has initiated and served as the Founding Chair of the International Conference on Food Properties (iCFP) and initiated the ICFP Mentoring Program for the young scientists and academics. He received the HortResearch Chairman’s Award, BRAP Award, CAMS Outstanding Researcher Award 2003, and SQU Distinction in Research Award 2008. In 2008, Professor Rahman has ranked among the top five Leading Scientists and Engineers of 57 OIC Member Countries in the Agroscience Discipline. Professor Rahman is an eminent scientist and academic in the area of food processing. He is recognized for his significant contribution to the basic and applied knowledge of food properties related to food structure, engineering properties and food stability. His total Google Scholar citations is more than 20,000 (h-index: 61) (16 November 2017), indicating high impact of his research in the international scientific community. Mohamed Al-Khusaibi Dr. Mohammed Al-Khusaibi, Assistant Professor at the Sultan Qaboos University, Sultanate of Oman. He has completed his B. Sc. (Food Science) in 2000 and M. Sc. (Food Science) in 2005 from Sultan Qaboos University. He has completed his Ph. D. degree in 2011 from the University of Reading, UK. Chapter 11 William E. Bornak William Bornak is a water treatment chemist specializing in water plant operations (ion exchange and RO), ultrapure water systems, and boiler steam cycle chemistry. He currently is Chief Technical Officer of Recirculation Technologies, LLC, having founded the company in 2000. The company was purchased in 2012 by a local Philadelphia investment group; Mr. Bornak has stayed on in a technical capacity. He developed and patented a new anion resin cleaning protocol and is extending the application to non-traditional resin uses. The resin lab, under his direction, has developed several tests to document the performance and financial gains from resin cleaning. Between 1996 and 2000 Bornak was principal consultant in Aqueous Solutions, a company dedicated to ultrapure water systems.

List of Contributors

731

Mr. Bornak worked at (then) Betz Laboratories in Trevose, PA for twenty years as ion exchange troubleshooter and manager of new hire chemical training. He gave several hundred seminars world-wide in water treatment. His first job out of school was at Rohm & Haas, then an independent company in Philadelphia, in R&D exploring new analytical methods for pesticide residues and for several years in the ion exchange division, specializing in the application of adsorbent resins to waste treatment problems. In addition to over fifty technical papers, Bornak has published a book on the practical operation of ion exchange demineralizers. He is an honors graduate from La Salle University in Philadelphia and did graduate work at the University of Pennsylvania. Chapter 12 Pei Xu Dr. Pei Xu is an Assistant Professor in the Department of Civil Engineering at New Mexico State University. Her research has focused on water and wastewater treatment engineering; treatment and beneficial use of produced water; desalination; membrane processes; potable and non-potable water reuse; water quality; and photocatalysis. She has authored and coauthored over 100 papers, book chapters, technical reports and conference presentations. She received her B.S in Environmental Engineering from Xi’an University of Architecture & Technology, M.S. in Water and Wastewater Engineering from Lanzhou Jiaotong University, China, and Ph.D. in Hydrosciences from Ecole Nationale du Génie Rural, des Eaux et des Forêts (ENGREF), France. Dr. Xu is a proposal reviewer for RPSEA/Department of Energy, Bureau of Reclamation, Civilian Research and Development Foundation (CRDF), Water Research Foundation (Formerly AwwaRF), NASA EPSCoR panel, and Qatar National Research Fund (QNRF). She is also a regular reviewer for scientific journals such as Environmental Science and Technology, Water Research, Journal of Membrane Science, Desalination, and Separation and Purification Technology, etc. Chapter 13 D. Yogi Goswami Prof. Yogi Goswami is the Distinguished University Professor and the Director of the Clean Energy Research Center at the University of South Florida.

732 List of Contributors He is the Editor-in-Chief of the “Solar Energy” journal, and “Progress in Solar Energy.” Dr. Goswami has published 21 books, 35 book chapters, and 390 refereed technical papers. He holds 21 patents, with some of the inventions such as the photocatalytic detoxification and disinfection systems becoming commercialized and are available worldwide. Products based on his inventions are helping allergy and asthma sufferers globally. His contributions in combined power/cooling cycles have impacted global research in a new class of thermodynamic cycles His inventions and innovations in solar energy has led to his being inducted into the Florida Inventors Hall of Fame in 2016. He has served as the President of the International Solar Energy Society (ISES), a Governor of ASME-International, Senior Vice President of ASME and President of the International Association for Solar Energy Education (IASEE). His work was named as one of the top 25 inventions for 2017 from Time Magazine. He won the 2017 Global Visionary Award from USF World. Dr. Goswami is the recipient of the prestigious Florida Inventors Hall of Fame 2016, Karl Böer Solar Energy Medal of Merit Award Univ. of Delaware; the Farrington Daniels Award from ISES (highest award of ISES); Frank Kreith Energy Award and medal from ASME (Highest Energy award of ASME); John Yellott Award for Solar Energy from ASME and the Charles Greely Abbott award (highest award of the American Solar Energy Society) and Hoyt Clark Hottel award of the ASES; and more than 50 awards and certificates from major engineering and scientific societies for his work in energy efficiency and renewable energy. Eydhah Almatrafi Eydhah Almatrafi is a faculty member of the college of engineering at the King Abdulaziz University in Saudi Arabia. He worked at ARAMCO and SABIC before winning a fellowship that changed his path from industry to academia through the King Abdulaziz University to complete his graduate studies in mechanical engineering in 2012. Eydhah is a PhD candidate at the University of South Florida. His Ph.D. research is focused on solar desalination, specifically, the analysis of thermal desalination assisted by vapor compression. He completed his M.Sc., in mechanical engineering at George Washington University in 2014 and earned his B.Sc., from King Fahd University of Petroleum and Minerals in 2010. He is a co-founder and board member of GCC-US Renewable Energy Association in Saudi Arabia that focuses on promoting and enabling

List of Contributors

733

renewable energy in GCC countries. He is a co-founder of the ABAAD foundation established in 2012 encouraging Saudi students who study in the US to start up a business. Chapter 14 Author Biography: Dr. José A. Carta is Professor of Mechanical Engineering at the University of Las Palmas de Gran Canaria (ULPGC), Canary Islands, Spain. Since the 1980s his research has centred on the field of renewable energies, with particular emphasis on wind energy. His activities as a researcher extend from the measurement and analysis of wind energy to the planning, modelling and development of applications of this renewable energy source. He currently heads the Group for the Research on Renewable Energy Systems (GRRES) at ULPGC. Numerous research studies that he has undertaken have focussed on the development of innovative wind-driven desalination systems. Author Biography: Dr. Pedro Cabrera is a research fellow in the Group for the Research on Renewable Energy Systems (GRRES) of the University of Las Palmas de Gran Canaria (ULPGC), Canary Islands, Spain. His current research work is centred on the application of wind energy in seawater desalination using reverse osmosis technologies, with special emphasis on employing Artificial Intelligence based control techniques to maximize exploitation of the available energy resource. His field of research interest also extends to the maximisation of renewable integration in stand-alone energy systems, flexible consumption, and the development of smart energy strategies. Author Biography: Dr. Jaime González Hernández has been a Professor at the University of Las Palmas de Gran Canaria (ULPGC), Canary Islands, Spain, in the Department of Electronic Engineering and Automation since the 1980s. He was also with Canterbury University, Christchurch, New Zealand, working on power system harmonics from 1992 to 1993. He is currently a member of the Group for the Research on Research on Renewable Energy Systems (GRRES) at ULPGC, undertaking research activities in the field of power electronics in relation to stand-alone electrical systems and microgrids. Chapter 15 Veera Gnaneswar Gude Veera Gnaneswar Gude is an Associate Professor of Environmental Engineering at Mississippi State University. He has over 15 years of academic,

734 List of Contributors research and industrial experience in desalination, wastewater treatment and biofuel scientific and technological areas. He received a BS degree in Chemical Engineering from Osmania University and worked for Du Pont Singapore for four years. He received a MS in Environmental Engineering from National University of Singapore and a PhD in Environmental Engineering from New Mexico State University under the direction of Prof. N. Nirmal Khandan for the research in low temperature thermal desalination. He is a licensed professional engineer and a board certified environmental engineer (also known as diplomate of environmental engineering) by American Academy of Environmental Engineers and Scientists (AAEES). Dr. Gude has published over 100 scientific research articles on desalination, water- wastewater treatment and biofuels research including 75 journal papers, 15 invited book chapters, 35 peer-reviewed conference proceedings papers, 20 technical reports, and 2 patents. He delivered 35 invited lectures including keynote lectures and over 150 scientific research and educational presentations. He organized many workshops on waterenergy-environment nexus topics at national and international conferences. He was the chair and board representative for clean energy and water division of ASES between 2011 and 2016. He serves on numerous scientific advisory boards and task committees across the world including ASCE-EWRI, ASEE, ASES, and AWWA and many scientific journals and research agencies. He received recognition for research and education from professional societies like ASCE, ASEE, ASES, and Chi Epsilon National Civil Engineering Honor Society. Chapter 16 Patrick V. Brady Dr. Pat Brady is a Senior Scientist at Sandia National Laboratories and has authored or co-authored several dozen peer-reviewed journal articles, books, book chapters, eighteen  patents, in the fields of water treatment, enhanced oil recovery, contaminant chemistry, and climate change. He joined Sandia in 1993 and also serves as adjunct Assistant Professor of Civil and Environmental Engineering at New Mexico Institute of Mining and Technology, Socorro, New Mexico. Education: Bachelors Geology, University of California –Berkeley (1984); Ph.D. - Geochemistry, Northwestern (1990); Post-doc - ETH-EAWAG, Dübendorf, Switzerland. Michael M. Hightower Mr. Hightower is a Research Professor at the University of New Mexico in the areas of water and energy research and development after serving 38

List of Contributors

735

years at Sandia National Laboratories in Albuquerque, New Mexico. Mike holds Bachelor’s and Master’s degrees in civil and environmental engineering from New Mexico State University and has 40 years of experience in space, weapons, and energy and natural resources research and engineering. His focus in the past two decades has been on the use of distributed and renewable energy and water treatment technologies to enhance economic development, global public health, and infrastructure and natural resource security and resiliency. Since 1992, Mike has supported the Departments of Interior and Energy and the Environmental Protection Agency in establishing science and technology evaluation programs and national roadmaps. These efforts have included improving innovative environmental technology cost and performance validation, addressing the environmental challenges of inland brackish water desalination, identifying major water and energy interdependencies, and identifying climate impacts on long-term, economic, and natural resource security, sustainability, and resiliency. These efforts have included the development of over 120 technical papers and reports for Congress, the Departments of State and Defense, the EPA, the National Science Foundation, the Electric Power Research Institute, the World Bank, and the World Energy Council. Mike has also supported the United States, Canada, Australia, the Middle East, and Asia in developing regional desalination and energy and water research programs and designing and developing national desalination research and testing facilities.

Index

Acid Mine Drainage, 482 Adsorption, 498–499, 500, 501 Advantages of freezing-melting, 447–448, 450–451, Aerodynamic braking, 581 AEROGEDESA project, 612–614, 617 Airgap Membrane Distillation (AGMD), 332, 334, 340 Angle of attack, 578–579 Antiscalant, 357 Antiscalants, 154, 161, 182 APASRENA Programme, 611 Aquaporin, 210, 224, 231, 232 Arsenic, Selective Removal, 487 Atmospheric emissions, 673 Bakken, 691 Beryllium, 692 Beta, 148–149 Beverage industry, 191 Biomimetic desalination, 714 Bitumen, 712 Blade, 575, 576 Boron, 657, 665, 667, 673, 692, 713 Boron, Selective Removal, 486 BRACKISH WATER, 3, 13–17, 29, 30, 33, 37, 38 Brine, 558, 660–662, 664, 667, 669, 674 Brine Recycle Mode, 84 Bubble column dehumidifier, 426–427 Bubble columns, 396, 409, 426–427, 428–430, 431–433

Calcium sulfate scale, 696 Capacitative Deionization (CDI), 500–502, 516 Capacitive deionization, 288 Capacitive deionization (CDI), 276, 277 Capacity factor, 647, 651–653, 676 Capital cost, 269 Capital costs, 676 Carbon dioxide, 307, 309–310 CARBON FOOTPRINT, 36, 39 Carbon nanotubes, 714 Carrier gases, 408 Caustic addition, 144 CDI treatment cystem, 502–503 Cell pair, 246 Cell thickness, 294 Cellulose acetate, 168 Cellulose triacetate membrane (CTA), 218, 219, 222, 223, 227, 228, 231 Chelants, 487 Chemical cleaning, 320, 322 Chemical desalination, 528 Chemical precipitation, 498 Cleaning-in-place (CIP), 264 CLIMATE CHANGE, 6, 7 Closed air (CA) cycle, 390 Closed water (CW) cycle, 390 Closed-air open-water cycle (CAOW), 532 Closed-air open-water water-heated cycle (CAOW), 399–401, 406

Jane Kucera (ed.) Desalination 2nd Edition, (737–745) © 2019 Scrivener Publishing LLC

737

738

Index

Closed-water open-air cycle (CWOA), 532 Coagulation-flocculation, 498 Cogeneration, 647, 656–657, 668, 670, 676, 678 Coil-free bubble column, 431–432 Co-ion, 249 Colloidal fouling, 181 Co-location, 649 Concentrate management, 714 Concentrated Solar Power (CSP), 553, 555, 556 Concentrating collectors, 556 Concentration Factor, 54, 140 Concentration mode, 190 Concentration polarization, 147–148, 149, 180, 253, 270, 336, 338, 354 Concentration polarization effects, 147 Concentration polarization factor see beta, CONSERVATION, 1, 4, 6, 9–11 Constant current, 272 Constant voltage, 266 Contamination, 498 Continuous diafiltration, 190 Continuous electrodeionization (CEDI), 276–278 Continuous Ion Exchange, 490 Control subsystem, 589–590 Controls, 304–306 Cooling, 647, 649, 651–652, 655, 662, 665, 673–674 Cost estimation, 268 Counter-ion, 249 Crystallization unit, 450, 452, 454–455, 457 Current efficiency, 301–302, 316 Current utilization, 259, 266, 268, 269 Cycles of concentration, 697 Dairy industry, 189–191 Dams, 685 Darcy's law, 158, 338

DC power, 304–305, 311–312, 316–317, 323 Deep well injection, 700 Degree of the hydrophobicity, 252 Dehumidifier, 396, 400, 409, 426–427, 428–430, 431–433, 435 Demineralization, 483 Desal Process, 480 Desalination, 526, 527, 529, 533, 541–542, 545, 550, 553, 554, 559, 568, 598, 600, 606, 609, 617, 620, 621, 627, 630, 187–189, 358, 364, 367 Desalination membrane processes, 246 Designer Waterflooding, 702 Desorption, 501 Dielectric exclusion, 171 Diffusion, 394, 426 Dilute and concentrate compartment, 246 Direct Contact Membrane Distillation (DCMD), 332, 334, 340 Direct solar desalination, 529–530 Direct use, 647, 651, 659, 669, 673, 676 Direct-contact freezing, 451–452, 459–460, 469 District heating, 651–652, 655 Donnan exclusion, 165, 171 Donnan steric pore model, 177–178 Doubly fed induction (DFIG) generator, 588 Draw solution, 222, 225, 226, 232, 233, Drive train, 579, 586 brake, 580–582 components, 580 configuration, 580–581 gearbox, 580, 582 shafts, 580582 Dual-Purpose Plant, 113, 117, 123 Dusty Gas Model (DGM), 339–340 ED process design, 262–264 Effective cell pair area, 266

Index Effectiveness model, 398, 407, 408, 409, 419 Eilat, Israel RO Installation, 488 El Paso Water Utilities, 693 Electodialysis, 499 Electrial subsystem, 584–585, 586 Electric pulse, 274 Electrical resistance, 251, 252 Electrochemical cycle stages, 501 Electrode materials, 303–304 Electrode reactions, 296–297, 303, 318 Electrodes, 500 Electrodialysis, 287–288, 298, 300–301, 315 Electro-dialysis, 545, 548 Electrodialysis reversal (EDR), 569, 609, 247 Electrolytes, 507 Electroneutrality, 173 Electroregeneration, 291 Electrosorption, 510, 514 Electrostatic bonds, 702–704 Energy costs, 659 Energy efficiency, 359 Energy requirement, 265–267 Energy-acquisition subsystem, 575–579 Enhanced oil recovery, 702 Enhanced transfer, 291, 295 Enthalpy, 395, 397, 402, 417 Enthalpy Pinch, 419 Enthalpy Pinch Model, 417–419 Entire load area, 591 Entrainment Factor, 94, 101 Entropy, 394, 396, 402, 410–411, 416, 425 Environmental considerations, 647 European Commission (DG XII) JOULE research programme, 622, 623 Eutectic, 451, 459–460 Faraday's Law, 300–301, 316 Faulkenmark Indicator, 3

739

Feed water, 647, 660–662, 665, 667–670, 675 Feed water quality, 504–506 Feed water requirements, 307 FGD wastewater, 232 Fixed speed wind turbine, 585–586 Flow spacers, 298 Flowback, 708, 709 Flow-through capacitors (FTC), 501 Fluid Technology Solutions (FTS), 213, 233 Fluorocarbon-type membranes, 250 Flux, 145–146, 334, 338, 341, 357 Flux decline, 181 Flux rate, 146 Footprint, 655, 669–671, 674 Fossil fuel, 526, 568, 590 Foulants, 270 Fouled membrane issues, 150 Fouling, 502, 149–152, 180 Fouling control, 180 Fouling index, 271 Fouling layer formation model, 272 Fouling minimization guidelines, 150 Fouling mitigation, 273, 274 Fouling of ion exchange membrane, 270 Fouling potential, 271–272 Fouling species, 149 Frac fluid chemistry, 704 Fracking wastewater, 232 Fresh Water Importation, 688 Gained Output Ratio, 91, 103, 111, 113 Gained Output Ratio (GOR), 359, 360, 391, 394, 403, 404, 420, 426, 435 Gaskets, 261 Gearboxes, types, 580 Generator, 584, 585–586, 587–589, 621, 630, 637 Geothermal energy, 647, 649–652, 654–657, 659, 661–663, 667–669, 671, 674, 677

740

Index

Geothermal source, 647, 649, 651–653, 656, 663, 670–672, 676, 678 GREENHOUSE GASES, 6, 14 GREY WATER, 17, 18 GROUND WATER, 10, 12, 14 Halite saturation, 707 Hardness, 144, 306–308, 319, 322 Heat loss, 342 Heat pump, 650, 655 Heat pump configuration, 541–542 Heat recovery factor, 360 Heat transfer, 336–337, 341–343, 355 Heavy metals, 498, 507, 657, 674 History of freezing-melting, 448 Hollow fiber, 213, 214, 222, 224–230 Homogenous and heterogeneous membranes, 249, 250 Homogenous solution model (HSDM), 179 Horizontal-axis wind turbine, 576 classification, 575 components, 574–575 Hub, 575, 576 Humidification-dehumidification, 657, 663, 668 Humidification-dehumidification desalination (HDH) cycle, 387, 390, 416, 422, 433 advantages of, 388 cycle classification, 390–391, 399 cycle components, 388–389, 395–396 effect of salinity, 433–437 performance parameters, 391–394 Humidifier, 395, 396, 400, 407, 409, 433, 435 Hybrid Plants, 131 Hydrate, 453, 456, 459–460, 466 Hydration Technology Innovation (HTI), 213, 219, 222–224, 226, 232, 233 Hydrogen gas, 303, 315 Hydrophilicity, 222, 223, 225, 231

Hydrophobicity, 225 Hydrostatic coupling, 617, 628 Ice crystallizaton, 450–452, 464 Ice separation, 450–451, 455, 457, 462 Impaired Waters, 692 Indirect solar desalination, 529–530, 533 Influent volume, 140, 142 Instrumentation, 304–306 Integrated ED process, 275, 279 Integrated Uses, 655 Internal concentration polarization (CP), 213, 220–223, Ion Exchange, 480, 498 Ion exchange capacity, 251, 252 Ion exchange membranes, 249–251, 287–289, 292, 299–300 Ion Exchange Metal Removal, 486 Ion exchange resin, 287–288, 290–301 Ion exchange resin, layered-bed, 294–296 Ion exchange resin, mixed-bed, 294–295 Ion exchange resin, separate-bed, 296–297 Ion Exchange Softening, 483 Ion Exchange vs. RO, 480 Ion product, 153 Irreversible thermodynamic model, 178–179 Janus type membrane, 351 K. Bailey Hutchison desalting plant, 692 Langelier Saturation Index (LSI), 160–161 Large-scale desalination system, 620–621, 635–636 Limiting current density, 259–260 Lithium, 698

Index Mass extraction, 417, 419, 420, 421, 426 Mass flow rate ratio, 400 Mass injection, 417, 419, 420, 421, 426 Mass transfer, 336–337, 355 resistance, 338, 340, 343 transport steps, 337, 338 Mass transfer resistance, 213, 219 Mass transport and feed boundary layer, 337–338 Mass transport and membrane pores, 338–341 Mechanical coupling, 617, 628 Mechanical power transmission subsystem, 579–583 Mechanical Vapor Compression, 712 Mechanical Vapour Compression (MVC), 569 Medium-scale desalination system, 620–621, 634 Membrane Capacitative Deionization (MCDI), 501 Membrane desalination, 647, 651, 660–661, 663, 667, 675–676 Membrane distillation, 657, 663, 668, 330, 340, 341 advantages of, 330 and renewable energy, 363–364, 365 applications, 331 challenges to, 330–331, 351, 355, 362, 364 configurations, 330, 332, 334 cost elements in, 361 energy and cost analyses, 358–364 future prospects, 367–369 history of, 334–336 hydrophobic porous membrane preparation, 344–345 innovations in, 365–367 membrane criterias, 343–344 module design, 351–356 process parameters, 356–358 working principle, 331–332 Membrane filtration transport mechanisms, 336

741

Membrane fouling, 180–182 Membrane fouling see fouling, Membrane wetting, 344, 345, 356 Merry-Go-Round Geometry, 487 Microfiltration, 344 Microfiltration (MF), 499 Microfiltration membrane, 158 Mixed Bed Polisher, 483 Modified fouling index, 158–159 Modified fouling index issues, 159 Modified Heat Capacity Rate Ratio (HCR), 397, 402, 409, 410, 416 Molecular diffusion, 340 Monovalent ion permselective membranes, 250 Multi-effect desalination (MED), 536–538, 541, 568 Multi-effect distillation, 657, 659, 664, 675–676 Multiple-Effect Distillation, 61 Multi-pole synchronous (MSG) generator, 587 Multi-stage air heated cycle, 407 Multi-stage bubble column dehumidifier, 428–430 Multi-stage flash (MSF), 568 Multistage freezing-melting, 449 Multi-Stage-Flash Evaporation, 80 Nacelle, 583, 584, 589, 590 Nanofiber, 224, 227, 230 Nanofiltration, 697, 703, 163, 165, 168–170, 180–181 applications, 182–193 ion rejection, 171–173 modeling separations, 177–180 organic solute rejection, 173–176 solute removal mechanisms, 171 Nanofiltration (NF), 499 Nanofiltration membrane, characteristics, 165, 166–167 classes, 165 Nanotechnology, 714

742

Index

Negative rejection, 173 Negatively charged group, 249 Niofouling, 181, 190 Non-concentrating collectors, 556 NORM, 692 Nuclear desalination, 568 Oasys Water, 213, 232 Off-design model see single-stage fixed area model, Off-grid wind energy system, 608 Ohm's Law, 302, 316 Oil recovery, 704 Oil sands, 700, 710, 712 On-Design model see effectiveness model, On-grid wind energy system, 600–602, 620 Open-air (OA) cycle, 390 Organic fouling, 181 Osmosis, 542 Osmotic pressure, 147, 210–212, 214, 215, 217, 220, 221, 228 Osmotic process, 213, 214, 221 Oxidants, 307, 310–311 Packed Bed, 492 Palo Verde Nuclear Power Plant, 698 Passive vaccum, 550–553 Performance Ratio, 112,133 Permeate Gap Membrane Distillation (PGMD), 352–353 Permeation rate, 357 Permselectivity, 251 Pharmaceutical industry, 192 Phase change desalination, 533–534 Phenomenological model see irreversible thermodynamic model, Photovoltaic (PV) solar energy, 278 Photovoltaic module, 279 Photovoltaics (PV) cell, 557–558, 559 Physical desalination, 528 Pinch Point, 395, 434

Pitch, 578, 579, 581, 587 Plate-and-frame, 291–292 Polarity reversal, 274 Polarization, 665 Pollution see contamination, Porosity, 215, 220–222, 224, Positively charged groups, 249 Powdered Resins, 493 Power coefficient, 590 Power curve, 591, 597, 578 Power transmission system, 582 Power-water cogeneration, 553 Practical membrane area, 267 Pressure balancing, 317–318 Pressure-shift nucleation, 455, 466 Principle of freezing-melting, 450 Process design parameters, 264 Produced water, 702 Product water, 659, 666–667, 669 Progressive layer crystallization, 460–461, 463 Rain cycle, 387 Record keeping, 319, 321 Recovery, 140, 312–313, 319 Recovery Rate, 54 Recovery ratio, 392 Rectifier, 304, 317 RECYCLE, 1, 9–11, 17 Recycling of reject stream, 310, 313 Refrigerant, 447–448, 451–456, 464–466 Reject concentration, 140 Rejection, 142–145, 148 Renewable energy, 526, 529, 568, 278 biomass, 568 geothermal, 568 ocean thermal, 568 solar, 568 tidal power, 568 wave, 568 wind, 568, 569 RENEWABLE ENERGY SOURCES (RES), 1, 35, 36, 39, 41, 42

Index Resin Polymers, 489 REUSE, 1, 4, 9–11, 14, 16, 17 Reverse Electrodialysis (RED), 276, 277 Reverse osmosis, 451, 465, 469–470, 164 Reverse osmosis (RO), 499, 568, 569 Reverse osmosis system recovery, 140–141 Reverse osmosis, rejection charateristics, 142–143 Reverse osmosis, system rating, 139, 163 RO flush, 308 Rotor, 576–578, 589 Rotor control, 578–579 active stall, 578, 579 full span pitch, 578–579 partial span pitch, 578, 579 passive pitch, 578, 579 passive stall, 578 Salinity feed, 433–437 Salinization, 685, 699 Salt passage, 142, 145, 148 Sanitization, 319, 323 Scaled membrane issues, 153 Scaling, 502, 152–154, 161, 182, 190 Scaling minimization guidelines, 152 Scaling species, 152 SCARCITY, 1, 4, 27 SEAWATER, 2, 3, 6–8, 11–18, 23, 24, 26, 30, 34, 35, 37, 38, 41 Seawater Desalination, 688 Seawater Desalination with an Autonomous Wind Energy System (SDAWES), 623, 624 Seawater greenhouse, 532–533 Seawater reverse osmosis (SWRO), 615, 631, 633 Second law efficiency, 436–437 Seismic activity, 700 Selective precipitation, 714 Semiconductors, 557 Shadow effect, 577 Shale, 705 Silica scale, 696, 697, 698

743

Silica scale inhibitors, 712 Silt density index (SDI), 146, 154–157 Silt density index test, 154–155 Silt density index test conditions, 156 Single-Purpose Plant, 113 Single-stage air heated cycle, 406–407 Single-Stage Evaporation, 52 Single-stage fixed area model, 408–410, 421 Size exclusion, 171, 178 Slickwater, 706 Small-scale desalination system, 608, 634 Sodium softening, 144, 161 Softening membrane, 168 Solar, 447–448, 468–470 Solar collectors, 535, 536, 553, 554, 556, 557 Solar desalination, 528, 529, 553, 554 Solar driven Heat pumps (HP), 541–542, 544 Solar driven humidificationdehumidification (HDH), 532–533 Solar driven multi-stage flash, 534–535 Solar driven Reverse osmosis (RO), 533, 542, 544–545 Solar energy, 527, 528, 529, 533 Solar flash, 550 Solar ponds, 555, 556–557, 558 Solar still, 530–531, 662 Solar stills, 389–390 Solar-driven Membrane distillation (MD), 548–550 Solute adsorbtion mechanisms, 175 Solute flux, 177 Solute rejection, 178 Spacer, 261, 262 Specific Electricity Consumption (SEC), 392 Specific Evaporator/Condenser Area, 73, 76, 90 Specific flux, 146–147

744

Index

Specific Primary Energy Consumption, 125, 126, 130 Specific Thermal Energy Demand, 63, 71, 77, 82, 83, 86 Spiral wound, 291–293 Stack design, 261–262 Steam Assisted GravityDrainage, 710 Steric hindrance see size exclusion, Stokes radius, 174 Strong Acid Cation Exchanger, 484 Strong Base Anion Exchanger, 484 Structural parameter, 221, 223, 224, 227 Sugar industry, 191 Supercooling, 450 Surface-force pore model (SFPM), 180 Sweep Gas Membrane Distillation (SGMD), 332, 334, 335 SWRO Reject as Regenerant, 485 Temperature polarization, 336, 342, 354 Temperature Polarization Coefficient (TPC), 342 Termal efficiency, 359, 362 Terminal Temperature Difference, 59 Terminal Temperature Difference (TTD), 406, 408 Textile industry, 192–193 Thermal conductivity, 343 Thermal desalination, 528, 647, 649, 659, 661–662, 669 Thermal efficiency, 389 Thermal imbalance, 397 Thermal separations, 246 Thermodynamic balancing, 419, 425 Thermoelectric, 648 Thermoelectric power generation, 696 Thick cell, 292 Thickness, 221, 223, 225 Thin cell, 292 Tight formations, 704, 705 Torque, 571, 578, 582 Tortuosity, 220, 221, 223, 224, Total Dissolved Solids, 658 Total organic carbon (TOC), 313

Transmission shafts, 580 Transport number, 251, 252 Treat to Need, 705 shale Trevi Systems, 213, 232, 233 Triple point, 464–465 Types of freezing-melting, 447, 451 Ultrafiltration (UF), 499 Ultrafiltration membrane, 159 Unconventional formations, 705 Uniform Particle Size (UPS) Resins, 493 Vaccum Membrane Distillation (VMD), 332, 334, 335, 340, 363 Vacuum-freezing, 464–465 Vapor-Compression, 464–465 Vapour Compressor, 98 Variable speed wind turbine, 586, 587–589 Varied pressure cycle, 407 VC mechanically driven, 104 VC thermally driven, 91, 97 Wash column, 447, 451–453, 457–460 Waste heat, 553, 362, 363, 365 Waste water, 498, 502 Waste water treatment, 183, 186–187, 192–193 WASTEWATER, 1, 10, 11, 14–17, 19, 30, 38, 42 Water Alliance, 629 Water Composition, 657 Water Production Cost (WPC), 359, 360, 362, 364 Water recovery ratio, 365 Water splitting, 287, 291, 295–296, 301, 308, 319 WATER STRESS, 1, 3–6 Water swelling, 251, 253 Water transport coefficient, 146 Water treatment, 182–183 Waterflooding, 704

Index Water-heated closed-air open-water cycle, 409, 417, 422 Weak Acid Cation Exchanger, 481, 485 Weak Base Anion Exchanger, 481 WELL WATER, 18 Wind desalination, 569, 600, 627 categories, 569 off-grid energy system, 608 on-grid energy system, 600–602, 620 system classification, 598–600 wind maching classification, 570–571 Wind energy characteristics, 590–591 Wind farm, 602, 603, 606, 634 Wind machine, classical application, 571 fast, 570–571, 572 horizontal-axis, 570–571, 575–579, 579–583 slow, 570–571, 608 vertical-axis, 570–571 Wind monitoring, 592, 593, 637–638 Wind power density (WPD), 594 Wind resource assessment, 591–598 Wind rose, 596 Wind turbine, 572, 573, 582, 586, 592, 599, 602–603, 623, 630 blade, 575, 576 classification, 574, 575 components, 574–575

745

control subsystem, 589 electrical subsystem, 584–585, 586 energy-acquisition subsystem, 575–579 fixed speed, 585–586 hub, 575, 576 mechanical power transmission subsystem, 579–583 power extraction influencers, 590 rotor, 576–578, 582, 583–584 rotor control, 578–579 support subsystem, 589–590 variable speed, 586, 587–589, 621 wind resource assessment, 591–598 yaw subsystem, 583–584 Wind turbine tower, 589, 590 Wind-diesel desalination, advantages, 633 drawbacks, 633–634 system classification, 630 system modes, 633 World Health Organization, 5, 6, 14 World Resources Institute, 3, 5 Yaw, 583–584, 589 Yaw subsystem, 583–584 Zero liquid discharge, 714 Zero liquid discharge (ZLD), 232, 331, 365 Zero-liquid discharge, 708