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Separation, extraction and concentration processes in the food, beverage and nutraceutical industries
© Woodhead Publishing Limited, 2010
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Related titles: Separation processes in the food and biotechnology industries (ISBN 978-1-85573-287-2) This book reviews methods and techniques for separating food components and products of the biotechnology industry. The introduction focuses on food composition and some of the conventional separation techniques. Subsequent chapters deal with each specific type or area of application individually and include information on the basic principles, industrial equipment available, commercial applications and an overview of research and development. Novel enzyme technology for food applications (ISBN 978-1-84569-132-5) The food industry is constantly seeking advanced technologies to produce valueadded, nutritionally-balanced products for consumers in a sustainable fashion. Since enzymes are so specific in their action, they are a useful biotechnological processing tool and by controlling the action of enzymes, innovative food ingredients and higher quality food products can be produced. Part one of Novel enzyme technology for food applications covers the principles of industrial enzyme technology, including methods to develop and tailor enzymes for food bioprocessing. Part two introduces the reader to novel applications of enzymes for the production of improved ingredients and food products. Food processing technology (Third edition) (ISBN 978-1-84569-216-2) The first edition of Food processing technology was quickly adopted as the standard text by many food science and technology courses. The publication of a completely revised and updated third edition consolidates the position of this textbook as the best single-volume introduction to food manufacturing technologies available. The third edition has been updated and extended to include the many developments that have taken place since the second edition was published. In particular, advances in microprocessor control of equipment, ‘minimal’ processing technologies, functional foods, developments in ‘active’ or ‘intelligent’ packaging, and storage and distribution logistics are described. Technologies that relate to cost savings, environmental improvement or enhanced product quality are highlighted. Additionally, sections in each chapter on the impact of processing on food-borne micro-organisms are included for the first time. Details of these and other Woodhead Publishing books can be obtained by: ∑ visiting our web site at www.woodheadpublishing.com ∑ contacting Customer Services (e-mail:
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© Woodhead Publishing Limited, 2010
iii Woodhead Publishing Series in Food Science, Technology and Nutrition: Number 202
Separation, extraction and concentration processes in the food, beverage and nutraceutical industries Edited by Syed S. H. Rizvi
Oxford
Cambridge
Philadelphia
New Delhi
© Woodhead Publishing Limited, 2010
iv Published by Woodhead Publishing Limited, Abington Hall, Granta Park, Great Abington, Cambridge CB21 6AH, UK www.woodheadpublishing.com Woodhead Publishing, 525 South 4th Street #241, Philadelphia, PA 19147, USA Woodhead Publishing India Private Limited, G-2, Vardaan House, 7/28 Ansari Road, Daryaganj, New Delhi – 110002, India www.woodheadpublishingindia.com First published 2010, Woodhead Publishing Limited © Woodhead Publishing Limited, 2010 The authors have asserted their moral rights. This book contains information obtained from authentic and highly regarded sources. Reprinted material is quoted with permission, and sources are indicated. Reasonable efforts have been made to publish reliable data and information, but the authors and the publisher cannot assume responsibility for the validity of all materials. Neither the authors nor the publisher, nor anyone else associated with this publication, shall be liable for any loss, damage or liability directly or indirectly caused or alleged to be caused by this book. Neither this book nor any part may be reproduced or transmitted in any form or by any means, electronic or mechanical, including photocopying, microfilming and recording, or by any information storage or retrieval system, without permission in writing from Woodhead Publishing Limited. The consent of Woodhead Publishing Limited does not extend to copying for general distribution, for promotion, for creating new works, or for resale. Specific permission must be obtained in writing from Woodhead Publishing Limited for such copying. Trademark notice: Product or corporate names may be trademarks or registered trademarks, and are used only for identification and explanation, without intent to infringe. British Library Cataloguing in Publication Data A catalogue record for this book is available from the British Library. ISBN 978-1-84569-645-0 (print) ISBN 978-0-85709-075-1 (online) ISSN 2042-8049 Woodhead Publishing Series in Food Science, Technology and Nutrition (print) ISSN 2042-8057 Woodhead Publishing Series in Food Science, Technology and Nutrition (online)
The publisher’s policy is to use permanent paper from mills that operate a sustainable forestry policy, and which has been manufactured from pulp which is processed using acid-free and elemental chlorine-free practices. Furthermore, the publisher ensures that the text paper and cover board used have met acceptable environmental accreditation standards. Typeset by Replika Press Pvt Ltd, India Printed by TJI Digital, Padstow, Cornwall, UK
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Contents
Contributor contact details..................................................................
xiii
Woodhead Publishing Series in Food Science, Technology and Nutrition............................................................................................... xvii Preface................................................................................................. xxvii Part I Developments in food and nutraceutical separation, extraction and concentration techniques 1 Principles of supercritical fluid extraction and applications in the food, beverage and nutraceutical industries................. Ž. Knez, M. Škerget and M. Knez Hrnčič, University of Maribor, Slovenia 1.1 Introduction....................................................................... 1.2 Thermodynamic fundamentals.......................................... 1.3 Cycle processes for extraction using supercritical fluids 1.4 Extraction of solids using SCF......................................... 1.5 Extraction of liquids using SCF........................................ 1.6 Conclusion......................................................................... 1.7 References......................................................................... 2 Principles of pressurized fluid extraction and environmental, food and agricultural applications................. C. Turner and M. Waldebäck, Uppsala University, Sweden 2.1 Introduction....................................................................... 2.2 Instrumentation and principles of pressurized fluid extraction . ........................................................................ 2.3 Applications of pressurized fluid extraction..................... 2.4 Future trends..................................................................... 2.5 Sources of further information and advice....................... 2.6 Conclusions ...................................................................... 2.7 References......................................................................... © Woodhead Publishing Limited, 2010
3 3 8 21 26 30 32 36 39 39 41 56 59 61 63 64
vi Contents 3 Principles of physically assisted extractions and applications in the food, beverage and nutraceutical industries..................................................................................... 71 E. Vorobiev, Compiègne University of Technology, France and F. Chemat, University of Avignon and Pays de Vaucluse, France 3.1 Introduction....................................................................... 71 3.2 Pulsed electric field-assisted extractions in the food industry.............................................................................. 72 3.3 Ohmic heating-assisted extractions in the food industry 83 3.4 Extraction assisted by high-voltage electrical discharges and applications in the food industry................................ 86 3.5 Ultrasound-assisted extraction (UAE) in the food industry.............................................................................. 90 3.6 Microwave-assisted extraction (MAE) in the food industry.............................................................................. 96 3.7 Combination of physical treatments for extraction in the food industry............................................................... 100 3.8 References......................................................................... 102 4 Advances in process chromatography and applications in the food, beverage and nutraceutical industries .................... M. Ottens and S. Chilamkurthi, Delft University of Technology, The Netherlands 4.1 Introduction....................................................................... 4.2 Basic principles of process chromatography.................... 4.3 Applications of process chromatography in the food, beverage and nutraceutical industries............................... 4.4 Recent developments in process chromatography............ 4.5 Process control in chromatography................................... 4.6 Future trends..................................................................... 4.7 Conclusions....................................................................... 4.8 Sources of further information and advice....................... 4.9 List of abbreviations......................................................... 4.10 References......................................................................... 5 Novel adsorbents and approaches for nutraceutical separation.................................................................................... B. W. Woonton, CSIRO Food and Nutritional Sciences, Australia and G. W. Smithers, Food Industry Consultant, Australia 5.1 Introduction....................................................................... 5.2 Molecular imprinted polymers and applications in the nutraceutical industry........................................................ 5.3 Organic monoliths and applications in the nutraceutical industry..............................................................................
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109 109 113 118 128 135 135 137 137 137 138 148
148 149 153
Contents vii 5.4 5.5 5.6 5.7 5.8 5.9
Stimuli-responsive resins and applications in the nutraceutical industry........................................................ Mesoporous molecular sieves and applications in the nutraceutical industry........................................................ Peptide affinity ligands and phage display methodology and applications in the nutraceutical industry................... Membrane adsorbers, membrane chromatography and applications in the nutraceutical industry......................... Conclusions and sources of further information and advice................................................................................ References.........................................................................
6 Advances in the effective application of membrane technologies in the food industry.............................................. M. Pinelo, G. Jonsson and A. S. Meyer, Technical University of Denmark, Denmark 6.1 Introduction....................................................................... 6.2 Theoretical fundamentals of membrane separation.......... 6.3 Membrane technology in the dairy industry..................... 6.4 Membrane technology in the fruit juice industry............. 6.5 Membrane technology for treatment of wastewater in the food industry............................................................... 6.6 New applications of membrane technology for the food industry: concentration and fractionation of saccharides 6.7 Future trends..................................................................... 6.8 References......................................................................... 7 Electrodialytic phenomena, associated electromembrane technologies and applications in the food, beverage and nutraceutical industries............................................................. L. Bazinet, A. Doyen and C. Roblet, Laval University, Canada 7.1 Introduction....................................................................... 7.2 Principles of electrodialytic phenomena and associated membrane technologies..................................................... 7.3 Applications of electrodialytic phenomena and associated membrane technologies................................... 7.4 Future trends .................................................................... 7.5 References......................................................................... 8 Principles of pervaporation for the recovery of aroma compounds and applications in the food and beverage industries..................................................................................... S. Sahin, Middle East Technical University, Turkey 8.1 Introduction....................................................................... 8.2 Principles of pervaporation............................................... 8.3 Transport mechanism in pervaporation for the recovery of aroma compounds......................................................... © Woodhead Publishing Limited, 2010
159 163 166 169 172 173 180 180 181 182 185 190 191 195 197
202 202 203 204 213 214
219 219 220 221
viii Contents 8.4 8.5 8.6 8.7
Selection of membranes for pervaporation in the recovery of aroma compounds.......................................... Recovery of aroma compounds by pervaporation and applications in the food and beverage industries.............. Sources of further information and future trends............. References.........................................................................
9 Advances in membrane-based concentration in the food and beverage industries: direct osmosis and membrane contactors ................................................................................... E. Drioli and A. Cassano, Institute on Membrane Technology, ITM-CNR, Italy 9.1 Introduction....................................................................... 9.2 Conventional technologies in the food and beverage industries........................................................................... 9.3 Direct osmosis and applications in the food and beverage industries............................................................ 9.4 Membrane contactors and applications in the food and beverage industries............................................................ 9.5 Conclusions....................................................................... 9.6 Nomenclature.................................................................... 9.7 References......................................................................... 10 Separation of value-added bioproducts by colloidal gas aphrons (CGA) flotation and applications in the recovery of value-added food products.................................................... P. Jauregi and M. Dermiki, The University of Reading, UK 10.1 Introduction ...................................................................... 10.2 Colloidal gas aphrons (CGA) properties........................... 10.3 Applications of CGA in the recovery of value-added food products..................................................................... 10.4 Feasibility of industrial application of CGA..................... 10.5 Future trends..................................................................... 10.6 Sources of further information and advice....................... 10.7 References......................................................................... 11 Membrane bioreactors and the production of food ingredients................................................................................... M.-P. Belleville, D. Paolucci-Jeanjean and G. M. Rios, European Institute of Membranes, France 11.1 Introduction....................................................................... 11.2 Membrane bioreactors for the production of food ingredients......................................................................... 11.3 Applications of membrane bioreactors in food industries. 11.4 Future trends..................................................................... 11.5 References.........................................................................
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227 230 239 240
244 244 245 248 250 275 275 278
284 284 285 293 307 308 309 310 314 314 315 322 331 331
Contents ix Part II Separation technologies in the processing of particular foods and nutraceuticals 12
Separation technologies in dairy and egg processing............. G. Gésan-Guiziou, INRA, France 12.1 Introduction....................................................................... 12.2 The dairy industry and composition of dairy products..... 12.3 Pretreatment of milk using separation techniques............ 12.4 Standardization and concentration of milk proteins in the dairy industry.............................................................. 12.5 Isolation of whole casein in the dairy industry................. 12.6 Separation techniques applied to whey and derivatives in the production of cheese............................................... 12.7 Fractionation of individual proteins and peptides in the dairy industry.................................................................... 12.8 Treatment of effluents and technical fluids in the dairy industry.............................................................................. 12.9 Conclusions and future trends in the dairy industry......... 12.10 The egg products industry and composition of egg products............................................................................. 12.11 Concentration and stabilization of egg white and whole egg..................................................................................... 12.12 Industrial extraction of egg-white proteins....................... 12.13 Industrial extraction of yolk components......................... 12.14 Conclusions and future trends in the egg-processing industry.............................................................................. 12.15 Sources of further information and advice....................... 12.16 References.........................................................................
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341 341 343 347 351 354 357 360 366 368 369 371 371 374 375 376 377
Separation technologies in the processing of fruit juices....... G. Vatai, Corvinus University of Budapest, Hungary 13.1 Introduction....................................................................... 13.2 Characteristics of foods/fluids in the fruit juice product sector................................................................................. 13.3 Designing separation processes to optimize product quality in the fruit juice product sector............................. 13.4 Production of fruit juice concentrate ............................... 13.5 References.........................................................................
381
1 4
396
Separation technologies in oilseed processing ........................ M. A. Williams, Anderson International Corp., USA 14.1 Introduction....................................................................... 14.2 Preparation for oilseed processing.................................... 14.3 Extrusion preparation for oilseed processing.................... 14.4 Mechanical pressing of oilseeds....................................... 14.5 Percolation solvent extraction in oilseed processing........ 14.6 Solvent recovery in oilseed processing............................. © Woodhead Publishing Limited, 2010
381 382 383 386 394
396 397 399 403 415 422
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14.7 14.8 14.9 14.10
Obtaining oil from fruit pulps........................................... Future trends..................................................................... Sources of further information and advice....................... References.........................................................................
424 425 427 428
1 5
Separation technologies in brewing.......................................... G. J. Freeman, Campden BRI, UK 15.1 Introduction....................................................................... 15.2 Characteristics of brewery products.................................. 15.3 Selection of technology and raw materials appropriate to brewery products........................................................... 15.4 Wort production in the brewhouse.................................... 15.5 Whirlpools and applications in brewing........................... 15.6 Yeast flocculation and applications in brewing................ 15.7 Beer fining agents ............................................................ 15.8 Filter aid filtration and applications in brewing................ 15.9 Regenerable and reusable filter aids and applications in brewing......................................................................... 15.10 Bulk beer filtration by membranes.................................... 15.11 Recovery of cleaning detergents in brewing..................... 15.12 Dissolved gas control by membrane technology.............. 15.13 Future trends..................................................................... 15.14 References.........................................................................
430
1 6 Methods for purification of dairy nutraceuticals.................... C. J. Fee, J. M. Billakanti and S. M. Saufi, University of Canterbury, New Zealand 16.1 Introduction ...................................................................... 16.2 Components of acidic whey protein ................................ 16.3 Purification technologies for acidic whey proteins . ........ 16.4 Basic proteins in the dairy nutraceutical industry............. 16.5 Purification technologies for basic whey proteins in the dairy nutraceutical industry............................................... 16.6 Immunoglobulins in the dairy nutraceutical industry....... 16.7 Purification technologies for immunoglobulins in the dairy nutraceutical industry............................................... 16.8 Future trends..................................................................... 16.9 References ........................................................................ 17 Methods of concentration and purification of omega-3 fatty acids . ......................................................................................... S. P. J. Namal Senanayake, Danisco USA, Inc., USA 17.1 Introduction....................................................................... 17.2 Urea adduction in the concentration and purification of omega-3 fatty acids........................................................... 17.3 Chromatographic methods for the concentration and purification of omega-3 fatty acids................................... © Woodhead Publishing Limited, 2010
430 431 432 433 434 435 436 437 441 443 446 446 447 448 450 450 451 454 462 463 470 471 473 474 483 483 484 486
Contents xi 17.4 Low-temperature fractional crystallization for the concentration and purification of omega-3 fatty acids...... 17.5 Supercritical-fluid extraction for the concentration and purification of omega-3 fatty acids................................... 17.6 Distillation methods for the concentration and purification of omega-3 fatty acids................................... 17.7 Enzymatic methods for the concentration and purification of omega-3 fatty acids................................... 17.8 Integrated methods for the concentration and purification of omega-3 fatty acids................................... 17.9 Conclusions....................................................................... 17.10 References......................................................................... 18 Extraction of natural antioxidants from plant foods . ........... E. Conde, A. Moure, H. Domínguez and J. C. Parajó, University of Vigo, Spain 18.1 Introduction....................................................................... 18.2 Antioxidant activity in food systems................................ 18.3 Natural compounds with antioxidant activity and major sources............................................................................... 18.4 Biological activities of natural antioxidants..................... 18.5 Extraction of natural antioxidants from plant foods and residues.............................................................................. 18.6 Integration of extraction processes and purification......... 18.7 Future trends..................................................................... 18.8 Sources of further information and advice....................... 18.9 Acknowledgements........................................................... 18.10 References......................................................................... 19 Fractionation of egg proteins and peptides for nutraceutical applications.......................................................... B. P. Chay Pak Ting, Y. Pouliot and S. F. Gauthier, Laval University, Canada and Y. Mine, University of Guelph, Canada 19.1 Introduction....................................................................... 19.2 Composition and physicochemical characteristics of egg proteins and applications in the nutraceutical industry.............................................................................. 19.3 Biological activities of egg proteins and peptides and applications in the nutraceutical industry......................... 19.4 Available technologies for the fractionation of egg proteins and peptides, and applications in the nutraceutical industry........................................................ 19.5 Conclusion and perspectives............................................. 19.6 References.........................................................................
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488 490 492 495 498 501 502 506 506 507 511 521 526 556 567 567 568 568 595
595 597 601 605 612 613
xii Contents 2 0 Supercritical-fluid extraction of lycopene from tomatoes...... J. Shi and S. Jun Xue, Agriculture and Agri-Food Canada, Canada, Y. Jiang, The Chinese Academy of Sciences, China and X. Ye, Zhejiang University, China 20.1 Introduction ...................................................................... 20.2 Supercritical-fluid extraction (SFE) of lycopene.............. 20.3 Factors affecting lycopene yield....................................... 20.4 Effects of pressure and temperature on the antioxidant activity of lycopene........................................................... 20.5 Effect of co-solvent and modifiers in lycopene extraction........................................................................... 20.6 Solubility of lycopene in supercritical fluids.................... 20.7 Conclusion and future trends ........................................... 20.8 References.........................................................................
619
619 622 623 628 631 634 639 640
Index..................................................................................................... 647
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Contributor contact details
Chapter 2
(* = main contact) Editor S. S. H. Rizvi Department of Food Science Cornell University 114B Stocking Hall Ithaca, NY 14853-7201 USA E-mail:
[email protected]
Chapter 1 Prof. Dr Ž. Knez*, Prof. Dr M. Škerget and M. Knez Hrnčič Faculty of Chemistry and Chemical Engineering University of Maribor Slomškov trg 15 2000 Maribor Slovenia
C. Turner* and M. Waldebäck Department of Physical and Analytical Chemistry Uppsala University P.O. Box 599 751 24 Uppsala Sweden E-mail:
[email protected]
and C. Turner Department of Chemistry Lund University P.O. Box 124 221 00 Lund Sweden E-mail:
[email protected]
E-mail:
[email protected]
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xiv Contributor contact details Chapter 3 E. Vorobiev* Groupe Technologies AgroIndustriels EA 4297 Département Génie des Procédés Université de Technologie de Compiègne B.P. 20529 60205 Compiègne France E-mail:
[email protected]
F. Chemat UMR 408, Sécurité et Qualité des Produits d’Origine Végétale Université d’Avignon et des Pays de Vaucluse INRA 84000 Avignon France E-mail:
[email protected]
Chapter 4 Marcel Ottens* and Sreekanth Chilamkurthi Delft University of Technology Department of Biotechnology Julianalaan 67 2628 BC Delft The Netherlands E-mail:
[email protected]
Chapter 5 B. W. Woonton* CSIRO Food and Nutritional Sciences 671 Sneydes Road (Private Bag 16) Werribee Victoria 3030 Australia E-mail:
[email protected]
G. W. Smithers Food Industry Consultant P.O. Box 158, Highett Melbourne Victoria 3190 Australia E-mail:
[email protected]
Chapter 6 M. Pinelo and A. S. Meyer* Department of Chemical and Biochemical Engineering Center for BioProcess Engineering Building 229 Technical University of Denmark DK-2800 Kgs. Lyngby Denmark E-mail:
[email protected]
G. Jonsson CAPEC Center for BioProcess Engineering Building 229 Technical University of Denmark DK-2800 Kgs. Lyngby Denmark Chapter 7 L. Bazinet*, A. Doyen and C. Roblet Institute of Nutraceuticals and Functional Foods (INAF) and Dairy Research Centre (STELA) Department of Food Sciences and Nutrition 2425 rue de l’agriculture, Pavillon Paul-Comtois Laval University Québec, QC, Canada G1V 0A6 E-mail:
[email protected]
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Contributor contact details xv Chapter 8
Chapter 12
S. Sahin Department of Food Engineering Middle East Technical University 06531, Ankara Turkey
Dr G. Gésan-Guiziou UMR1253 Science et Technologie du Lait et de l’œuf INRA – Agrocampus Ouest 65 rue de Saint Brieuc 35042 Rennes cedex France
E-mail:
[email protected]
E-mail: genevieve.gesan-guiziou@ rennes.inra.fr
Chapter 9 Enrico Drioli* and Alfredo Cassano Institute on Membrane Technology, ITM-CNR c/o University of Calabria via P. Bucci, 17/C I-87030 Rende (CS) Italy E-mail:
[email protected];e.drioli@ unical.it;
[email protected]
Chapter 10
Chapter 13 Prof. G. Vatai Corvinus University of Budapest Faculty of Food Science Department of Food Engineering 1114 Budapest Menesi út 44 Hungary E-mail:
[email protected]
Dr Paula Jauregi* and Dr Maria Dermiki Department of Food and Nutritional Sciences The University of Reading Whiteknights, P.O. Box 226 Reading RG6 6AP UK E-mail:
[email protected]
Chapter 14 M. A. Williams Anderson International Corp. 6200 Harvard Avenue Cleveland, OH 44105 USA E-mail:
[email protected]
Chapter 11 Dr M.-P. Belleville*, Dr D. Paolucci-Jeanjean and Prof. G. M. Rios Institut Européen des Membranes UMR 5635 CC 047 – UM2 Place E. Bataillon F 34095 Montpellier cedex 05 France E-mail: marie-pierre.belleville@iemm. univ-montp2.fr
Chapter 15 G. J. Freeman Campden BRI Centenary Hall Coopers Hill Road Nutfield Redhill Surrey RH1 4HY UK E-mail:
[email protected]
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xvi Contributor contact details Chapter 16
Chapter 19
C. J. Fee*, J. M. Billakanti and S. M. Saufi Biomolecular Interaction Centre Department of Chemical and Process Engineering University of Canterbury Private Bag 4800 Christchurch 8020 New Zealand
Bertrand P. Chay Pak Ting, Yves Pouliot*, Sylvie F. Gauthier Département de Sciences des Aliments et de Nutrition and Institute of Nutraceuticals and Functional Foods (INAF) 2425 rue de l’agriculture, Pavillon Paul-Comtois Laval University Québec, QC Canada G1V 0A6
E-mail:
[email protected]
E-mail:
[email protected]
Chapter 17 S. P. J. Namal Senanayake Danisco USA, Inc. Four New Century Parkway New Century, KS 66031 USA E-mail:
[email protected]
Chapter 18 E. Conde, A. Moure, H. Domínguez* and J. C. Parajó Department of Chemical Engineering Faculty of Sciences University of Vigo Campus Ourense Spain E-mail:
[email protected]
Y. Mine Department of Food Science University of Guelph Guelph, ON Canada N1G 2W1 Chapter 20 J. Shi* and S. Jun Xue Guelph Food Research Center Agriculture and Agri-Food Canada Ontario Canada N1G 5C9 E-mail:
[email protected]
Y. Jiang South China Botanical Garden The Chinese Academy of Sciences Guangzhou 510650 China X. Ye Department of Food Science and Nutrition School of Biosystems Engineering and Food Science Zhejiang University Zhejiang 310029 China
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xvii
Woodhead Publishing Series in Food Science, Technology and Nutrition
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xx Woodhead Publishing Series in Food Science, Technology and Nutrition 71 Safety and quality issues in fish processing Edited by H. A. Bremner 72 Minimal processing technologies in the food industries Edited by T. Ohlsson and N. Bengtsson 73 Fruit and vegetable processing: improving quality Edited by W. Jongen 74 The nutrition handbook for food processors Edited by C. J. K. Henry and C. Chapman 75 Colour in food: improving quality Edited by D MacDougall 76 Meat processing: improving quality Edited by J. P. Kerry, J. F. Kerry and D. A. Ledward 77 Microbiological risk assessment in food processing Edited by M. Brown and M. Stringer 78 Performance functional foods Edited by D. Watson 79 Functional dairy products Volume 1 Edited by T. Mattila-Sandholm and M. Saarela 80 Taints and off-flavours in foods Edited by B. Baigrie 81 Yeasts in food Edited by T. Boekhout and V. Robert 82 Phytochemical functional foods Edited by I. T. Johnson and G. Williamson 83 Novel food packaging techniques Edited by R. Ahvenainen 84 Detecting pathogens in food Edited by T. A. McMeekin 85 Natural antimicrobials for the minimal processing of foods Edited by S. Roller 86 Texture in food Volume 1: semi-solid foods Edited by B. M. McKenna 87 Dairy processing: improving quality Edited by G Smit 88 Hygiene in food processing: principles and practice Edited by H. L. M. Lelieveld, M. A. Mostert, B. White and J. Holah 89 Rapid and on-line instrumentation for food quality assurance Edited by I. Tothill 90 Sausage manufacture: principles and practice E. Essien 91 Environmentally-friendly food processing Edited by B. Mattsson and U. Sonesson 92 Bread making: improving quality Edited by S. P. Cauvain 93 Food preservation techniques Edited by P. Zeuthen and L. BøghSørensen 94 Food authenticity and traceability Edited by M. Lees 95 Analytical methods for food additives R. Wood, L. Foster, A. Damant and P. Key 96 Handbook of herbs and spices Volume 2 Edited by K. V. Peter 97 Texture in food Volume 2: solid foods Edited by D. Kilcast 98 Proteins in food processing Edited by R. Yada 99 Detecting foreign bodies in food Edited by M. Edwards 100 Understanding and measuring the shelf-life of food Edited by R. Steele © Woodhead Publishing Limited, 2010
Woodhead Publishing Series in Food Science, Technology and Nutrition xxi 101 Poultry meat processing and quality Edited by G. Mead 102 Functional foods, ageing and degenerative disease Edited by C. Remacle and B. Reusens 103 Mycotoxins in food: detection and control Edited by N. Magan and M. Olsen 104 Improving the thermal processing of foods Edited by P. Richardson 105 Pesticide, veterinary and other residues in food Edited by D. Watson 106 Starch in food: structure, functions and applications Edited by A-C Eliasson 107 Functional foods, cardiovascular disease and diabetes Edited by A. Arnoldi 108 Brewing: science and practice D. E. Briggs, P. A. Brookes, R. Stevens and C. A. Boulton 109 Using cereal science and technology for the benefit of consumers: proceedings of the 12th International ICC Cereal and Bread Congress, 24–26th May, 2004, Harrogate, UK Edited by S. P. Cauvain, L. S. Young and S. Salmon 110 Improving the safety of fresh meat Edited by J. Sofos 111 Understanding pathogen behaviour in food: virulence, stress response and resistance Edited by M. Griffiths 112 The microwave processing of foods Edited by H. Schubert and M. Regier 113 Food safety control in the poultry industry Edited by G. Mead 114 Improving the safety of fresh fruit and vegetables Edited by W. Jongen 115 Food, diet and obesity Edited by D. Mela 116 Handbook of hygiene control in the food industry Edited by H. L. M. Lelieveld, M. A. Mostert and J. Holah 117 Detecting allergens in food Edited by S. Koppelman and S. Hefle 118 Improving the fat content of foods Edited by C. Williams and J. Buttriss 119 Improving traceability in food processing and distribution Edited by I. Smith and A. Furness 120 Flavour in food Edited by A. Voilley and P. Etievant 121 The Chorleywood bread process S. P. Cauvain and L. S. Young 122 Food spoilage microorganisms Edited by C. de W. Blackburn 123 Emerging foodborne pathogens Edited by Y. Motarjemi and M. Adams 124 Benders’ dictionary of nutrition and food technology Eighth edition D. A. Bender 125 Optimising sweet taste in foods Edited by W. J. Spillane 126 Brewing: new technologies Edited by C. Bamforth 127 Handbook of herbs and spices Volume 3 Edited by K. V. Peter
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xxii Woodhead Publishing Series in Food Science, Technology and Nutrition 128 Lawrie’s meat science Seventh edition R. A. Lawrie in collaboration with D. A. Ledward 129 Modifying lipids for use in food Edited by F. Gunstone 130 Meat products handbook: practical science and technology G. Feiner 131 Food consumption and disease risk: consumer-pathogen interactions Edited by M. Potter 132 Acrylamide and other hazardous compounds in heat-treated foods Edited by K. Skog and J. Alexander 133 Managing allergens in food Edited by C. Mills, H. Wichers and K. Hoffman-Sommergruber 134 Microbiological analysis of red meat, poultry and eggs Edited by G. Mead 135 Maximising the value of marine by-products Edited by F. Shahidi 136 Chemical migration and food contact materials Edited by K. Barnes, R. Sinclair and D. Watson 137 Understanding consumers of food products Edited by L. Frewer and H. van Trijp 138 Reducing salt in foods: practical strategies Edited by D. Kilcast and F. Angus 139 Modelling microorganisms in food Edited by S. Brul, S. Van Gerwen and M. Zwietering 140 Tamime and Robinson’s Yoghurt: science and technology Third edition A. Y. Tamime and R. K. Robinson 141 Handbook of waste management and co-product recovery in food processing Volume 1 Edited by K. W. Waldron 142 Improving the flavour of cheese Edited by B. Weimer 143 Novel food ingredients for weight control Edited by C. J. K. Henry 144 Consumer-led food product development Edited by H. MacFie 145 Functional dairy products Volume 2 Edited by M. Saarela 146 Modifying flavour in food Edited by A. J. Taylor and J. Hort 147 Cheese problems solved Edited by P. L. H. McSweeney 148 Handbook of organic food safety and quality Edited by J. Cooper, C. Leifert and U. Niggli 149 Understanding and controlling the microstructure of complex foods Edited by D. J. McClements 150 Novel enzyme technology for food applications Edited by R. Rastall 151 Food preservation by pulsed electric fields: from research to application Edited by H. L. M. Lelieveld and S. W. H. de Haan 152 Technology of functional cereal products Edited by B. R. Hamaker 153 Case studies in food product development Edited by M. Earle and R. Earle
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xxiv Woodhead Publishing Series in Food Science, Technology and Nutrition 178 New technologies in aquaculture: improving production efficiency, quality and environmental management Edited by G. Burnell and G. Allan 179 More baking problems solved S. P. Cauvain and L. S. Young 180 Soft drink and fruit juice problems solved P. Ashurst and R. Hargitt 181 Biofilms in the food and beverage industries Edited by P. M. Fratamico, B. A. Annous and N. W. Gunther 182 Dairy-derived ingredients: food and neutraceutical uses Edited by M. Corredig 183 Handbook of waste management and co-product recovery in food processing Volume 2 Edited by K. W. Waldron 184 Innovations in food labelling Edited by J. Albert 185 Delivering performance in food supply chains Edited by C. Mena and G. Stevens 186 Chemical deterioration and physical instability of food and beverages Edited by L. H. Skibsted, J. Risbo and M. L. Andersen 187 Managing wine quality Volume 1: viticulture and wine quality Edited by A.G. Reynolds 188 Improving the safety and quality of milk Volume 1: milk production and processing Edited by M. Griffiths 189 Improving the safety and quality of milk Volume 2: improving quality in milk products Edited by M. Griffiths 190 Cereal grains: assessing and managing quality Edited by C. Wrigley and I. Batey 191 Sensory analysis for food and beverage quality control: a practical guide Edited by D. Kilcast 192 Managing wine quality Volume 2: oenology and wine quality Edited by A. G. Reynolds 193 Winemaking problems solved Edited by C. E. Butzke 194 Environmental assessment and management in the food industry Edited by U. Sonesson, J. Berlin and F. Ziegler 195 Consumer-driven innovation in food and personal care products Edited by S. R. Jaeger and H. MacFie 196 Tracing pathogens in the food chain Edited by S. Brul, P. M. Fratamico and T. A. McMeekin 197 Case studies in novel food processing technologies: innovations in processing, packaging, and predictive modelling Edited by C. Doona, K Kustin and F. Feeherry 198 Freeze-drying of pharmaceutical and food products T-C. Hua, B-L. Liu and H. Zhang 199 Oxidation in foods and beverages and antioxidant applications Volume 1: understanding mechanisms of oxidation and antioxidant activity Edited by E. A. Decker, R. J. Elias and D. J. McClements 200 Oxidation in foods and beverages and antioxidant applications
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201 202 203 204 205 206 207 208 209 210 211 212 213 214 215 216 217 218 219 220 221 222
Volume 2: management in different industry sectors Edited by E. A. Decker, R. J. Elias and D. J. McClements Protective cultures, antimicrobial metabolites and bacteriophages for food and beverage biopreservation Edited by C. Lacroix Separation, extraction and concentration processes in the food, beverage and nutraceutical industries Edited by S. S. H. Rizvi Determining mycotoxins and mycotoxigenic fungi in food and feed Edited by S. De Saeger Developing children’s food products Edited by D. Kilcast and F. Angus Functional foods: concept to profit Second edition Edited by M. Saarela Postharvest biology and technology of tropical and subtropical fruits Volume 1 Edited by E. M. Yahia Postharvest biology and technology of tropical and subtropical fruits Volume 2 Edited by E. M. Yahia Postharvest biology and technology of tropical and subtropical fruits Volume 3 Edited by E. M. Yahia Postharvest biology and technology of tropical and subtropical fruits Volume 4 Edited by E. M. Yahia Food and beverage stability and shelf-life Edited by D. Kilcast and P. Subramaniam Processed Meats: improving safety, nutrition and quality Edited by J. P. Kerry and J. F. Kerry Food chain integrity: a holistic approach to food traceability, authenticity, safety and bioterrorism prevention Edited by J. Hoorfar, K. Jordan, F. Butler and R. Prugger Improving the safety and quality of eggs and egg products Volume 1 Edited by Y. Nys, M. Bain and F. Van Immerseel Improving the safety and quality of eggs and egg products Volume 2 Edited by Y. Nys, M. Bain and F. Van Immerseel Feed and fodder contamination: effects on livestock and food safety Edited by J. Fink-Gremmels Hygiene in the design, construction and renovation of food processing factories Edited by H. L. M. Lelieveld and J. Holah Technology of biscuits, crackers and cookies Fourth edition Edited by D. Manley Nanotechnology in the food, beverage and nutraceutical industries Edited by Q. Huang Rice quality K. R. Bhattacharya Meat, poultry and seafood packaging Edited by J. P. Kerry Reducing saturated fats in foods Edited by G. Talbot Handbook of food proteins Edited by G. O. Phillips and P. A. Williams
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xxvi
xxvii
Preface
Separation, extraction and concentration of desirable components from their natural matrices are essential unit operations in the preparation of key ingredients for use in the food, pharmaceutical and chemical industries. These processes often account for somewhere between 50% and 70% of the product cost and are facing new challenges. Ever-tightening regulations on the use of organic solvents, environmental issues and process safety have accelerated the development of a variety of new technologies which are clean and efficient, and do not cause degradation of the products; these technologies therefore enhance the productivity and global competitiveness of the industries concerned. This book aims to provide a comprehensive overview of the most important technologies of interest for the production of high-value compounds. Based on the multidisciplinary expertise of 45 contributors from institutions with strong programs in separation, extraction and concentration processes, the book is organized in 20 peer-reviewed chapters, divided into two parts. Part I describes the latest advances in separation, extraction and concentration techniques, including supercritical fluid extraction, process chromatography and membrane technologies. It also reviews emerging techniques of particular interest, such as pervaporation and pressurized liquid extraction. Part II then focuses on advances in separation technologies and their applications in various sectors of the food, beverage and nutraceutical industries. Areas covered include dairy and egg processing, oilseed extraction and brewing. This part of the book discusses the characteristics of different foods and fluids, how food constituents are affected by separation processes and how separation processes can be designed and operated to optimize end product quality. This volume collectively provides valuable and timely information on the latest developments in the field. With its team of experienced international contributors, Separation, extraction and concentration processes in the food, beverage and nutraceutical industries is an important reference source for experienced professionals concerned with the development and optimization of these processes. It © Woodhead Publishing Limited, 2010
xxviii Preface is hoped that newcomers to this exciting and emerging field will also find valuable information in this book. I wish to thank the authors for their patience and support during the review and preparation of the manuscripts. I also wish to thank Sarah Whitworth of Woodhead Publishing who gave useful advice during the initial planning stages of the book. Syed S. H. Rizvi
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Principles of supercritical fluid extraction and applications 3
1 Principles of supercritical fluid extraction and applications in the food, beverage and nutraceutical industries Ž. Knez, M. Škerget and M. Knez Hrn�i�, University of Maribor, Slovenia Abstract: The thermodynamic fundamentals of supercritical fluid extraction (SFE) are described and the environmental, health and safety benefits of using supercritical fluids are explored. Several hundred industrial-scale SFE plants are in operation worldwide for extraction of plant materials, such as hop constituents, decaffeination of tea and coffee, and separation of lecithin from oil, all high-pressure processes. Smaller industrial units are used for extraction of spices in the food industry and for natural substances used in cosmetics. The design of such an extraction plant is described. The unique thermodynamic and fluid dynamic properties of dense gases are also applied in integrated extractions and in in situ formulations, such as impregnation of solid particles, formation of solid powder emulsions, and particle coatings. Key words: extraction, supercritical fluids, dense gases, high pressure, thermodynamics, food industry.
1.1 Introduction The design of new products with special characteristics or of new processes that are environmentally friendly and have an impact on sustainability, present a great challenge to chemical engineers. Within the human environment, pressures range from 0.25 bar at the top of the highest mountain, up to 1000 bar at the bottom of the deepest ocean. Because human beings live on the surface of the globe, the first technologies for the production of various substances took place at atmospheric pressure. In the early 20th century, demand for new products like ammonia shifted the technological processes towards high pressure. Industrial high-pressure processes operate at ranges from about 50 bar (in particle formation processes) to over 200 © Woodhead Publishing Limited, 2010
4 Separation, extraction and concentration processes kbar (conversion of graphite to diamonds). High pressure is a relatively new tool and in several processes it has resulted in completely new products with special characteristics. Many of these new processes are environmentally friendly, low cost and sustainable. The advantages of using supercritical fluids (SCF) as solvents in chemical synthesis offer environmental benefits, health and safety benefits and chemical benefits (Jessop and Leitner, 1999). The environmental benefits of most SCFs in industrial processes result from their replacement of far more environmentally damaging conventional organic solvents. The low energy consumption of the process is a further environmental benefit. Health and safety benefits include the fact that the most important SCFs (SC CO2 and SC H2O) are non-carcinogenic, non-toxic, non-mutagenic, non-flammable and thermodynamically stable. One of the major process benefits is derived from the thermophysical properties of SCFs: high diffusivity, low viscosity and the density and dielectric constant of SCF, which can be fine tuned by changes in operating pressure and/or temperature. The motivation for using high pressure in a wide range of technologies and processes is based on chemical, physicochemical, physicobiochemical, physicohydrodynamic and physicohydraulic effects (Bertuco and Vetter, 2001). The extraction of hop constituents and the decaffeination of tea and coffee are the largest scale processes and are mostly performed on an industrial scale. Several industrial plants also extract spices for the food industry and natural substances for use in cosmetics. The advantages of using supercritical fluids for the isolation of natural products have been well described (Marr and Gamse, 2000) and include solvent-free products, low temperature, and no byproducts. One of the most important advantages of using supercritical fluids is the selective extraction of components or the fractionation of total extracts, which is made possible by use of different gases for isolation or fractionation of components and/or by changing the process parameters. A limitation on the further application of high-pressure technology for obtaining extracts is the relatively high price of the product compared with that of those produced conventionally. The legal restrictions on solvent residues and solvents (in products for human use) and the isolation or fractionation of special components from total extracts, in combination with different formulation (controlled release for example) and sterilisation processes, encourage the use of dense gases in extraction applications. There are fewer industrial units involved in using supercritical fluids for the separation of components from liquid mixtures. Some laboratoryscale studies involve extraction systems using liquid/supercritical fluid. Some data on binary systems liquid/SCF has been produced, but there are fewer data on liquid/liquid/supercritical fluid systems. As in all extraction processes, including the supercritical extraction of solid and liquid mixtures, the solubility of a single component or a mixture of components in SCF is the basic data for the design of separation processes. The components
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Principles of supercritical fluid extraction and applications 5 or mixture of compounds to be extracted must be soluble in SCF/dense gas. As is known from thermodynamics, the solubility of compounds in SCF/ dense gases depends on the density of the SCF/dense gas, which depends on the pressure and temperature of the SCF. Another very important parameter influencing the solubility of compounds in SCF is its dielectric constant, which is influenced by the temperature and/or pressure of the SCF. The general flow sheet of the extraction process is presented in Fig. 1.1 and some industrial scale units are shown in Figs 1.2 to 1.4. In one extraction stage, the solubility of a compound or mixture of compounds has to be high whereas in another stage, the solubility of a compound in SCF has to be low. Therefore, the phase equilibrium data is the most important factor in the design of the operating pressure and temperature of SCF in an extraction plant. Thus, the theoretical amount of SCF necessary for separating compounds from a solid or a liquid mixture may be calculated. The design of process parameters has a very important influence on the investment costs of high pressure plants and subsequently, for the economics of the process. In addition to the solubility data of solute in SCF, mass transfer also has a large influence on the economics of the extraction process. Mass transfer models usually describe extraction yield versus extraction time, but a better presentation for the design of extraction apparatuses is yield versus mass of SC solvent/mass of solid material (S/F). Cascade operation is used in industrial-scale plants to increase the economy of the solid–SC solvent extraction process. To date, there has been some research on the continuous operation of plants for the extraction of solids with SC solvent (Eggers, 1996), but currently, no application for
Extractor
PEXT
PS
TEXT
TS
Solvent tank
Extract
Fig. 1.1 General flow sheet of SCF extraction plant. Subscripts EXT and S represent the extractor and separator respectively. © Woodhead Publishing Limited, 2010
6 Separation, extraction and concentration processes
Fig. 1.2 High-pressure extraction unit (700 bar). (Photo: courtesy of Uhde HPT, Hagen, Germany.)
Fig. 1.3 High-pressure extractor solids during manufacturing. (Photo: courtesy of Uhde HPT, Hagen, Germany.)
the continuous feed of solids has been applied on an industrial scale. In industrial scale operations, extractors are usually combined in series. By means of the cyclic operation of a battery of extractors, a quasi-continuous solid flow may be achieved. Such a mode of operation results in extremely high extraction yields because pure solvent is in contact with pre-extracted material, thus loading the solvent to maximum capacity.
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Principles of supercritical fluid extraction and applications 7
Fig. 1.4 High-pressure extraction unit: closure of extractor. (Photo: courtesy of Uhde HPT, Hagen, Germany.)
1
2
3
Solvent tank
4
Extract
Fig. 1.5 Cascade of extractors for extraction of solid materials.
One of the major advantages of SFE processes is the fractionation of extracts. Multi-step separation may be performed by use of several separators by decreasing the solvent power. Decreased solvent power may be achieved by various methods, as described in sections 1.3 and 1.4.
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8 Separation, extraction and concentration processes
Extractor
PEXT
PS1
PS2
PS3
TEXT
TS1
TS2
TS3
Fraction 1
Fraction 2
Solvent tank
Fraction 3
Fig. 1.6 Scheme of multi-step separation.
1.2 Thermodynamic fundamentals Extraction with dense gases is, as with solvent extraction and leaching, a separation process based on the solubility of compounds in phases present in the system. The driving force for mass and heat transfer is the difference from the equilibrium state under given conditions. High-pressure phase behaviour near critical points can be complex, even in simple binary mixtures, especially when the components differ in molecular size, shape, structure and polarity. Complex phase behaviour can be readily interpreted by P–T and P–x projections of P–T–x space diagrams. By using the phase rule (equation [1.1]), the geometrical limits for presentation of multiphase regions in the phase diagram are determined.
F = c + 2 – p
[1.1]
where F is the number of independent variables, c is the number of components and p is the number of phases. The phase equilibrium of compounds (1, 2, 3, ..., N) distributed in phases (a, b, g, ..., p) present in the system can be defined in terms of the fugacity (f) by the following equation:
fia = fib = fig = …… = fip
[1.2]
where i = 1, 2, 3, ..., N. Only binary systems will be discussed in detail in this chapter. The phase equilibria of ternary and multicomponent mixtures is discussed by Brunner (1994) and McHugh and Krukonis (1986), and Sadus (1992). 1.2.1 Solid–supercritical fluid equilibrium Phase diagrams In solid–SCF systems where the normal melting temperature of the solid is higher than the critical temperature (Tc) of the SCF, two possible types of © Woodhead Publishing Limited, 2010
Principles of supercritical fluid extraction and applications 9 phase behaviour exist. The simplest is presented in Fig. 1.7a and is typical of mixtures in which the components are chemically similar. The critical mixture curve runs continuously between the critical points of both components of the mixture. The solid–liquid–gas (SLG) line is continuous and begins at the normal melting point of the heavy component, moves toward lower temperatures as the pressure is increased, and ends at a temperature usually well below the critical temperature of the lighter component. The melting point of the pure solid normally increases with an increase in the hydrostatic pressure. However, in the presence of dense gas, the melting point of the solid decreases as the pressure increases owing to the increasing solubility of gas in the solid. The second type of solid–SCF phase behaviour (Fig. 1.7b) is typical for systems in which the solid and the SCF differ considerably in molecular size, shape and/or polarity and can be interpreted as type III fluid-phase behaviour (de Loos, 2006) according to the classification of van Konynenburg and Scott (1980). In this type of system, the light gas is not very soluble in the heavy liquid, even at high pressures. Therefore, the melting-point depression of the solid is relatively small. The SLG curve is no longer continuous; three phase SLG equilibria are presented by two branches of the SLG line in P–T diagram. The high-temperature branch of the SLG line starts at the normal melting point of the solid and intersects with the critical-mixture curve at the upper critical end point (UCEP). The low-temperature branch of the SLG line intersects with the critical-mixture curve at the lower critical end point (LCEP). At these two points, the liquid and gas phases merge into a single fluid phase in the presence of excess solid. Only solid–gas equilibria exist between these two branches of the SLG line. Possible phase behaviour for type III systems with interference of the solid phase is presented in detail by de Loos (2006). The course of the high-temperature branch of the three-phase SLG line of a binary system depends on the solubility of the gas in the liquid phase. SL
UCEP
1
C2
LG
C1
SLG
2
Pressure (P)
Pressure (P)
SL C1
LG
LG
C2
LCEP SLG
2
1 SLG
TP
TP
Temperature (T) (a)
Temperature (T) (b)
Fig. 1.7 Solid–SCF equilibrium: P–T projection of phase diagram for similar (a) and asymmetrical (b) binary systems. C, critical point; TP, triple point; L, liquid; G, gas; S, solid; UCEP, upper critical end point; LCEP, lower critical end point. Dashed curves are critical lines, lines denoted as 1 and 2 are the vapour pressure curves of the two components. © Woodhead Publishing Limited, 2010
10 Separation, extraction and concentration processes If a compressed gas is dissolved in the melting of a heavy component, two opposite temperature effects occur as given by equation [1.3] (Arons and Diepen, 1963):
fus Ê ∂T ˆ ˆ Ê dT = Á ˜ dP + Á ∂T ˜ dxA = TDVfus d + Ë ∂x A ¯ P Ë ∂P¯ x D A
2 A
fus
d
A
[1.3]
These two effects are that the increase of hydrostatic pressure increases the RT melting temperature of the heavy component, and Pthe dissolved gas x in the D H x H heavy component decreases the melting temperature. TheDhigher solubility of gas in the melted heavy component results in a larger melting-point depression. Four characteristic shapes of the SLG equilibrium lines in the P–T projection were observed experimentally for asymmetric binary systems of compressed gases and non-volatile components (Arons and Diepen, 1963; Loos, 2006; Tuminello et al., 1995; Weidner et al., 1997) (Fig. 1.8), with: ∑ negative dP/dT slope, where the effect of gas solubility predominates; ∑ positive dP/dT slope, where the effect of pressure predominates; ∑ temperature minimum, where both effects are competing; and ∑ temperature maximum and a temperature minimum. The last type of SLG line is a very rare phenomenon and can be explained by the higher solubility of the supercritical gas in the solid phase than in the liquid phase (de Loos, 2006). A SLG curve with a temperature maximum and a temperature minimum was reported for the systems CO2 + polyethylene glycol (Weidner et al., 1997), CO2 + tripalmitin (O’Connell et al., 2003), and
LG
UCEP
Pressure (P)
SLG 2
SLG 1 C2 SLG 4
SLG 3
1
TP
Temperature (T)
Fig. 1.8 Characteristic shapes of the SLG equilibrium lines.
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Principles of supercritical fluid extraction and applications 11 CO2-tristearin (O’Connell et al., 2003). However, in most studies of these equilibria, it is assumed that the supercritical gas is insoluble in the solid phase of the non-volatile component (de Loos, 2006). The course of the SLG line is dependent on the gas and the chemical structure of the compound, i.e. the type and position of the functional groups. As an example, the melting point of vitamin K3 under the pressure of various gases is presented in Fig. 1.9 (Knez and Škerget, 2001) and the various paths of the SLG line can be observed. In the presence of CO2 and dimethyl ether, the negative slope dP/dT can be observed and the meltingpoint depression of vitamin K3 is highest under the pressure of dimethyl ether. The melting-point depression of vitamin K3 is less pronounced in the presence of propane, the SLG curve having a minimum at 94.9 °C and 39 bar. Under the pressure of inert gas (nitrogen and argon), the SLG curve has a positive dP/dT slope owing to the low solubility of gas in vitamin K3. Another example, which illustrates that isomers may have a different type of SLG line in the presence of a specific gas, is the vanillin–gas system. In Fig. 1.10, SLG phase lines for binary systems of vanillin (V) and o-vanillin (o-V) with fluorinated hydrocarbons (R23, R134a, R236fa) and CO2 are presented. For vanillin with –OH group in the para position, the meltingpoint depression in CO2 and fluorinated hydrocarbons is generally lower as for vanillin with –OH group in the ortho position. Thermodynamic modelling For a binary solid–SCF two-phase system at equilibrium, the fugacity of the solute in the solid phase is equal to that in the supercritical phase:
fiS (P, T , x ) = fiG (P, T , y)
[1.4]
350
Propane Dimethyl ether
300
CO2 250
N2
P (bar)
Ar 200 150 100 50 0 50
70
90
T (°C)
110
130
150
Fig. 1.9 SLG phase equilibria for binary system K3 – gas (CO2, propane, dimethyl ether, argon or nitrogen) (Knez and Škerget, 2001).
© Woodhead Publishing Limited, 2010
12 Separation, extraction and concentration processes 400 V - R23
CHO
350
V - R134a
300
V - R236fa OH
V - CO2 P (bar)
250
OCH3
o-V - R23 CHO
o-V - R134a
200
OH
o-V - R236fa 150
OCH3
o-V - CO2
100 50 0
0
10
20
30
40 50 T (°C)
60
70
80
90
Fig. 1.10 SLG phase lines for binary systems of vanillin (V) and o-vanillin (o-V) with dense gases.
By denoting the gas as component 1 and the non-volatile compound as component 2, and assuming that the solubility of the gas in the solid phase is negligible (x2 = 1), the fugacity of the solid in the solid phase is equal to the fugacity of pure solid (Prausnitz et al., 1986):
f2S (P, T , x ) = f2S,pure (P, T )
[1.5]
The fugacity of a pure compound is:
RT ln
f2S,pure (P,T ) = P
Ú0
RT ln
f2S,pure (P,T ) = P
Ú0
P
P2S
RT ˆ Ê ÁË v2 – P ˜¯ dP RT ˆ Ê ÁË v2 – P ˜¯ dP +
[1.6] P
ÚP
S 2
Ê S RT ˆ ÁË v2 – P ˜¯ dP [1.7]
where P2s is the sublimation pressure at the system temperature and v2S is the molar volume of the pure solid. The first term on the right side is the fugacity of the saturated vapour, which is equal to the fugacity of the saturated solid phase. The second term is the correction owing to the compression of the solid phase to pressure P.
RT ln
f2S,pure (P,T ) fS = RT ln 2S + P P2
P
ÚP
S 2
v2S dP S P – RT ln P2 P
© Woodhead Publishing Limited, 2010
[1.8]
Principles of supercritical fluid extraction and applications 13 By rearranging and inserting the expression for the fugacity coefficient fS j 2S = 2S : P2
Ê f2S,pure (P, T ) = P2Sj 2S exp Á 1 Ë RT
P
ÚP
S 2
ˆ v2S dP˜ ¯
[1.9]
where j2S is the fugacity coefficient at T and P2S. The fugacity of the gas phase is:
f2G (P, T , y) = j 2G y2 P
[1.10]
j2G
is the fugacity coefficient of solid component 2 in the mixture. where By inserting equations [1.9] and [1.10] into [1.5] and [1.4], the expression for solubility of solids in the gas phase is obtained:
P Ê ˆ P2Sj 2S expÁ 1 Ú S v2S dP˜ Ë RT P2 ¯ y2 = G Pj 2
[1.11]
Equation [1.11] can also be expressed in the form:
y2 =
P2S E P
[1.12]
where
P Ê ˆ j 2S expÁ 1 Ú S v2S dP˜ Ë RT P2 ¯ E∫ G j2
[1.13]
The enhancement factor E is the correction of the ideal-gas expression that is valid only at low pressures and contains three terms: j2S, which takes into account the non-ideality of the pure saturated vapour. For low sublimation pressure of the solid P2S, j2S almost equals unity. ∑ an exponential term called the Poynting correction, which gives the effect of pressure on the fugacity of the pure solid. It is small at low pressures but may become larger at high pressures or at low temperatures. ∑ j2G, the gas-phase fugacity coefficient in the high-pressure gas mixture. This term is the most important, because it is much lower than 1 and can produce very large enhancement factors (103 or higher). ∑
Assuming that the molar volume of the pure solid (v 2S) at the system temperature is pressure independent, the Poyning correction takes the simple form:
© Woodhead Publishing Limited, 2010
14 Separation, extraction and concentration processes
Ê vS (P – P2S )ˆ j 2S expÁ 2 ˜¯ RT Ë E= G j2
[1.14]
j2G can be calculated from an equation of state using following equation:
lnj i = 1 RT
•
ÚV
ÈÊ ∂P ˆ ˘ – RT ˙ dV – ln PV ÍÁ ˜ nT RT V ˙ ÍÎË ∂ni ¯ T ,V , n j ˚
[1.15]
where ni is the molar number of species i in the mixture, V is the total volume, v is the molar volume, and R is the gas constant. To calculate fugacity coefficients, the equations of state (EOS) are commonly applied in engineering practice because they can be used to predict the thermodynamic properties of fluids and describe the phase behaviour of mixtures over a wide range of temperature and pressure. Detailed reviews on the thermodynamic models applied for predicting phase behaviour and modelling aspects in supercritical fluid mixtures have been presented elsewhere (Anderko, 1990, 2000; Dohrn, 1994; Tuminello et al., 1995). Cubic EOS in combination with mixing rules are currently the most widely used models for the calculation of solubilities of components in SCF (Tables 1.1 and 1.2). Table 1.1 Cubic equations of state General form
Van der Waals (VDW: b = 0, c = 0) Redlich–Kwong (RK: c = 0)
Redlich–Kwong–Soave (RKS: c = 0) Peng–Robinson (PR: c = b)
a P = RT – v – b v(v + b ) + c(v – b )
[1.16]
a P = RT – 2c v–b v
[1.17]
acTc1/2 P = RT – v – b v(v + b )T 1/2
[1.18]
a P = RT – v – b v(v + b )
[1.19]
a P = RT – v – b v(v + b ) + b(v – b )
Where: a = aca
ac = Wa
[1.20]
[1.21]
RT Pc 2
2 c
[1.22]
a = [1 + c(1 – Tr1/2)]2
[1.23]
2
c = A 0 + B 0w + C 0w
[1.24]
RTc Pc
[1.25]
b = Wb
© Woodhead Publishing Limited, 2010
Principles of supercritical fluid extraction and applications 15 Table 1.2 Constants in cubic equations of state Konstanta
VDW
RK
RKS
PR
Wa Wb A0 B0 C0
0.42188 0.125 – – –
0.42478 0.08664 – – –
0.42747 0.08664 0.48 1.574 –0.176
0.45724 0.0778 0.37464 1.54226 –0.26992
In the equations in Tables 1.1 and 1.2, the parameters a and b reflect the contribution of attractive forces and molecular volume, respectively, and w, the acentric factor, is a measure of the acentric nature of intermolecular forces. In order to extend the use of a pure-fluid EOS to mixtures, it is assumed that the EOS for the mixture is the same as for a hypothetical pure fluid (Prausnitz et al., 1986) and that the characteristic constants a and b are dependent on composition. Usually, van der Waals one-fluid mixing rules with one or two adjustable parameters are used:
a = ∑∑ yi y j aij i j
b = ∑∑ yi y j bij i j
[1.26]
[1.27]
where a accounts for interactions between the species in the mixture and b accounts for the excluded volume of the species of the mixture. The cross coefficients aij and bij are related to the corresponding pure-component parameters by the following combining rules:
aij = aii a jj (1 – kij )
bij = 1 (bii + b jj )(1 – cij ) 2
[1.28] [1.29]
where kij and cij are the binary interaction and size parameters, respectively. When cij is set to zero, the co-volume b is expressed by a linear mixing rule, equation [1.29] reduces to:
b = ∑ yi bi i
[1.30]
In most practical applications, a linear mixing rule is used for the co-volume b (Anderko, 1990). However, many authors use a quadratic mixing rule for b in analogy with the mixing rule for a. In this way, a second binary parameter is introduced, which is usually found to be useful in correlating gas–liquid equilibria in mixtures with components of very different size (Anderko, 1990). Similarly, it usually improves results for solid–supercritical fluid equilibria © Woodhead Publishing Limited, 2010
16 Separation, extraction and concentration processes because of the large size differences in such systems (Anderko, 1990). Generally, kij and cij are of the same order of magnitude (Dohrn, 1994) and both are expected to have an absolute value of much less than 1 (McHugh et al., 1986). They can both be positive or negative. A negative value of kij usually indicates that specific chemical interactions, such as hydrogen bonding, are present in the mixture (McHugh et al., 1986). However, it is less apparent how to interpret a negative value for cij. By calculating partial derivatives of cubic EOS, inserting derivatives in equation [1.15] and after integrating the expression for the fugacity coefficient of the component, i is obtained. The expression for the fugacity coefficient of component i by using the Peng–Robinson (PR) EOS with the above defined van der Waals mixing rules is:
lnj i =
bN Ê Pv ˆ – 1˜ b ÁË RT ¯
Ê N ˆ 2∑y j aij Á b ˜ P(v – b ) j a – ln – – N˜ Á b˜ RT 2 2 RTb Á a ÁË ˜¯
¥l
[1.31]
where:
N
N
N –1 N
k
j
j =1i = j +1
bN = 2∑yk bik – ∑ y 2j b jj – 2 ∑ ∑ yi yi –j
ij
and for a binary system: v + (1 + 2)b y b ln bN = 2yi bij +v 2+y(1 yi2bbii – y 2j b jj – 2yi y j bi j b jj– – 2)
[1.32]
[1.33]
ij in the above equations The critical properties and acentric factors needed can usually be estimated by the use of group contribution methods, molar volume can be determined experimentally e.g. with a pycnometer, whereas the binary interaction and size parameters, kij and cij, respectively, are obtained by fitting the equation of state to measured phase-equilibrium data. The form of the mixing rules that extend the use of EOS developed for pure fluids to mixtures is, as reported (Anderko, 1990; Škerget et al., 2002), more important than the particular P–V–T relationship embodied in EOS. The shortcoming of the van der Waals one-fluid mixing rules is that they are only applicable to mixtures that exhibit relatively moderate solution nonidealities. Further problems, which persist when modelling phase equilibria in SCF mixtures with the use of cubic equations of state where the systems are highly non-ideal owing to the high pressures involved, are described by Prausnitz et al. (1986) and Škerget et al. (2002):
(a) the interaction parameters are often found to be temperature dependent; © Woodhead Publishing Limited, 2010
Principles of supercritical fluid extraction and applications 17 (b) reliable values for the necessary physical parameters are not always available; and (c) the equation does not fit the data equally well at all temperatures and pressures, and especially in the vicinity of the critical point, the deviation from the model is high. There have been a number of attempts to extend the range of EOS. Most equations retain the van der Waals separation of repulsive and attractive terms, but have introduced some modifications to either the attractive or repulsive term, or to both. In this way, a number of empirical and theoretical models have been developed (Sandler and Orbey, 2000, Sadus, 1992), which allow extrapolation and prediction over wide ranges of temperature and pressure and can describe greater degrees of non-ideality. Empirical approaches to eliminating the shortcomings of the van der Waals one-fluid model for a cubic EOS have been done to provide additional composition dependence or density dependence, by adding parameters to the combining rule for the parameter a, generally leaving the mixing rule for the parameter b unchanged (Sadus, 1992, Sandler and Orbey, 2000, Škerget et al., 2002). Some examples are the composition dependent combining rules of Adachi and Sugie (1986), Panagiotopoulos et al. (1986), Sandoval et al. (1989) and Schwartzentruber and Renon (1989 a,b). The density-dependent mixing rules have been revived by Anderko (1990) and Danner and Gupte (1986). A more empirical way was adopted by Luedecke and Prausnitz (1985). However, there are several problems associated with these multi-parameter combining rules which limit their use (Anderko, 1990; Škerget et al., 2002). The common difficulty is that they do not result in the correct treatment of ternary and multicomponent mixtures (de Loos, 2006). An alternative method is based on the combination of the equations of state with activitycoefficient models (Huron and Vidal (1979) and Wong and Sandler mixing rules (Wong and Sandler, 1992), which are usually fairly reliable for the prediction of multicomponent phase equilibria from binary data, except for the most strongly non-ideal systems. In binary three-phase SLG systems at equilibrium, the parameters needed in the PR EOS to model the SLG line are determined by solving the following equations:
f 2G(T, P, y2) = f 2L(T, P, x2)
[1.34]
f 2S (T,
[1.35]
f 1G(T, P, y1) = f 1L(T, P, x1)
P) =
f 2G(T,
P, y2)
[1.36]
Equation [1.35] transforms to equation [1.11] and equations [1.36] and [1.34] become:
y1j1G = x1j1L
[1.37]
y2j2G
[1.38]
=
x2j2L © Woodhead Publishing Limited, 2010
18 Separation, extraction and concentration processes Calculation of the pure solute fugacity The fugacity of a solute in a solid phase cannot be directly calculated by a conventional EOS, but it may be calculated by means of a sub-cooled liquid reference state. An equation developed by Prausnitz et al. (1986) is used to express the fugacity of the sub-cooled liquid at temperature T in terms of measurable thermodynamic properties: ln
f2L,pure DH 2fus Ê Tt,2 ˆ = ÁË T – 1¯ S,pure RT t,2 f2
Dcp,2 Tt,2 R T
Dcp,2 R
Tt,2 T [1.39]
where Tt,2 is the triple-point temperature, and DH 2fus the enthalpy of fusion Ê ˆ for component 2 at temperature T. ˜ – – 1˜ + ln ÁË ¯ ¯ Usually, it is sufficient to consider only the first term in equation [1.39], especially, if T and Tt are not far apart, and to neglect the contribution of cp. Furthermore, because there is little difference between the triple-point temperature and the normal melting temperature, it is common to substitute Tt for its normal melting point Tfus,i. The simplified equation thus obtained, valid at the triple point pressure Pt of the solute, is: f2L,pure DH 2fus Ê Tfus,2 = –1 [1.40] ˆ f2S,pure RTfus,2 Ê T ÁË ˜¯ To take into account the effect of pressure, the fugacities may be written as: ln
f2S,pure (P, T ) = f2S,pure (Pt , T ) exp
Ê 1 RT
ÊÊ 1 f2L,pure (P, T ) = f2L,pure (Pt , T ) expÁ Ë RT
P
ÚP v2S d P t
ˆ
P
ÚP v2L d P˜¯
[1.42]
t
P
ˆ ˜¯ (v – v ) dP˘ ˙˚
and the fugacity ratio is:
Ê ÁË f2S,pure (P, T ) f2S,pure (Pt , T ) = L,p exp È 1 L,pure ÍÎRT f2 (P, T ) f2 (P , T )
[1.41]
ÚP
t
By inserting equation [1.40] pureinto equation [1.43]:
S 2
L 2
[1.43]
t
f2S,pure (P, T )
and by integrating:
=
È f2L,pure (P, T )exp Í
1 RT Î
P
ÚP (v2S – v2L )dP + t
D
fus 2
Ê Á1 – fus,2 Ë
fus,2 ˆ ˘
˜¯ ˙ ˚
[1.44]
P
È(vS – v2L )(P – Pt f2S,pure (P, T ) = f2L,pure (P, T ) exp Í 2 RT Î –
)
© Woodhead Publishing Limited, 2010
+
T H RT T DH 2fus Ê Tfus,2 ˆ ˘ 1 RTfus,2 ÁË T ˜¯ ˙˚ –
[1.45]
Principles of supercritical fluid extraction and applications 19 1.2.2 Liquid–SCF equilibrium Phase diagrams According to the classification proposed by van Konynenburg and Scott (1980), six main types of fluid-phase behaviour in binary mixtures are distinguished by their critical properties. The corresponding P–T projections are shown in Fig. 1.11. Similar behaviour is observed for type I and type II, where a continuous LG critical line, which connects the critical points of the pure components, is observed. However, the difference between both types is observed because type II liquid mixtures are not miscible in all proportions and exhibit a LLG three phase line at low temperatures. The LLG line ends at the UCEP, where two liquid phases merge into one liquid phase. From this point, the LL critical line rises rapidly to higher pressures. In type III phase behaviour, the LG critical line is not a continuous line connecting critical points of the pure components, instead it has two branches; one going from the critical point of the component with the higher critical temperature, initially in the direction of the critical point of the other component and then rising sharply with pressure, whereas the other line goes from the critical point of the other component to the UCEP. Various sub-classes of type III behaviour have been observed, depending on the course of the main critical line. The critical line starting from C2 with a positive slope, indicates the existence of the so-called gas–gas equilibria Type I
Type II
1 Pressure (P)
C2
LG
C1
LL
Type III
1
2
Type V LG
C2 LL
C2
C1
LG
UCEP
LCEP
UCEP LLG
2
LCEP
1 LLG
UCEP
LL
2
1
C2
LG
C1
LG
LLG
2
LLG
Type VI LG
C1
1
LG
UCEP LLG
Type IV
C2
LG C1
1
2
C2
LG
C1
UCEP LLG
2
LCEP
Temperature (T)
Fig. 1.11 Classification of the phase behaviour of binary fluid systems. C, critical point; L, liquid; G, gas; UCEP, upper critical end point; LCEP, lower critical end point. Dashed curves are critical lines, lines denoted as 1 and 2 are the vapour pressure curves of the two components (Brunner, 1994; Prausnitz et al., 1986).
© Woodhead Publishing Limited, 2010
20 Separation, extraction and concentration processes (Sadus, 1992). The two phases are in equilibrium at a temperature higher than the critical temperature of either pure component. The three-phase LLG line is also observed near the vapour pressure line of the most volatile compound. In type IV phase behaviour, the critical line between the critical points of pure compounds is interrupted by a three-phase LLG line. One branch of the critical line goes from the critical point of the component with the higher critical temperature to the LCEP, whereas the other branch goes, as in type III, from C1 to the UCEP. Type V is similar to type IV except that there is also an LL critical line at low temperatures which ends at another UCEP. For this type of mixture, there are therefore two regions of limited liquid miscibility at lower and higher temperatures and pressures. In type VI there are two critical curves: a continuous LG critical line, which connects the critical points of the pure compounds, and an LL critical curve, which connects the UCEP and LCEP located at either end of threephase LLG line. Thermodynamic modelling of high-pressure vapour–liquid (V–L) equilibria Consideration of binary liquid in equilibrium with a gas phase: the equilibrium equation for a binary liquid–gas system at temperature T and pressure P is:
f 1G(T, P, y1) = f 1L(T, P, x1)
[1.45]
f 2G(T, P, y2) = f 2L(T, P, x2)
[1.46]
where xi is the mole fraction of component i in the liquid phase and yi is the mole fraction of component i in the gas phase. The fugacity in each phase can be written as:
f Li = xi jLi P
[1.47]
f Gi = yi jGi P
[1.48]
where jGi and jLi are the fugacity coefficients of i in the gas and liquid phases, respectively.
y1j1V = x1j1L
[1.49]
y2jV2 = x2 j2L
[1.50]
Fugacity coefficients for each component in the liquid and gas phase can be calculated by using equation [1.15].
© Woodhead Publishing Limited, 2010
Principles of supercritical fluid extraction and applications 21
1.3 Cycle processes for extraction using supercritical fluids 1.3.1 Solvent cycle A typical high-pressure extraction process basically comprises a separation stage for the feedstock (extractor and separator) and a regeneration stage for the solvent. In the separation stage, the components to be extracted become concentrated in the gas and are then precipitated in the separator by applying suitable methods (Fig. 1.12). The gas must be removed from the extract and the raffinate and cleaned for reuse in the extraction process. If a solvent mixture is applied in the extraction process, the composition of the mixture must be adjusted before reuse. The solvent recovery can be performed in various ways and the procedure chosen depends on the nature of substances, the scale of the process unit and the operating conditions (Brunner, 1994). The extract is separated from the solvent by changing the operating conditions in such a way as to reduce the solvent power of the SCF. This can be achieved by various means (Brunner, 1994): ∑
Isenthalpic throttling (expansion) to subcritical conditions (changing pressure and temperature), ∑ Changing the temperature and maintaining supercritical conditions, or cooling down to subcritical conditions, ∑ Employing an additional mass separating agent (absorbing medium, membrane adsorbents), whilst maintaining supercritical conditions for the solvent. Separation stage
Material to be extracted
Regeneration stage
Gas
Solvent–solute mixture
Extractor
Separator
Gas reservoir
Gas Raffinate
Gas
Solvent (compressed gas)
Extract (solute)
Fig. 1.12 Solvent cycle in extraction process using SCF. © Woodhead Publishing Limited, 2010
22 Separation, extraction and concentration processes The appropriate method and conditions in the separator are chosen according to the phase equilibrium data. In general, a change in temperature will not be effective in cleaning the solvent sufficiently for reuse in the extraction process, but a change in temperature can be applied in addition if total regeneration of the solvent is not necessary (Brunner, 1994). The advantage of using an absorbent or adsorbent for separation of the extract and the solvent is that the separation can be performed without significantly reducing the pressure. Operating at constant pressure reduces energy consumption. The disadvantage is that separation of the extract and the absorbent/adsorbent have to be performed subsequently, and this may be difficult and must be taken into account when designing the process and calculating its operating costs. The solvent in the subcritical (liquid) or supercritical state may be driven either by a compressor or by a pump. Solvent cycles with corresponding changes of conditions of state can be best presented in T–S diagrams and are used for calculating the heat balance of each step of the process. Pump process In Fig. 1.13 the temperature–entropy (T–S) diagram is schematically presented with areas of homogeneous liquid, gas, SCF and two-phase liquid–gas region (L–G). The dotted line separates the area of SCF from the gas and liquid state area and presents no phase border. The part of the line on the right of the critical point (CP) lies on the critical isobar (73 bar for CO 2). The extraction process is performed under constant conditions (3 in Fig. 1.13). The SCF phase following the extraction leaves the extraction unit and the dissolved substance is subsequently separated from the solvent by changing the pressure and temperature to reduce the solvent power of the SCF (isenthalpic H6
P5
P4
H5
Temperature (T)
P3 H4 H3 H2 2
H1 L
Pc
SCF
3
P2 P1
CP
G 5
1
6
L–G
4
Entropy (S)
Fig. 1.13 Gas extraction process in T–S diagram: solvent circuit in the pump mode.
© Woodhead Publishing Limited, 2010
Principles of supercritical fluid extraction and applications 23 throttling 3–4). The two-phase region of the solvent is obtained and the substance dissolved in the solvent precipitates to form a separate phase that can be removed. The gaseous part of the solvent is cooled and condensed to a liquid state (6) and then sub-cooled (1). To remove the substances that remain dissolved, the liquid solvent can be evaporated (6–5), subsequently liquefied (5–6) and sub-cooled (1). The sub-cooled liquid is then pumped to the operating pressure (1–2) and heated to the operating temperature (2–3). The area enclosed by the solvent circuit in the T–S diagram represents the thermodynamic work needed for the process of cycling the SCF. The lower the pressure drop needed for separating the extract and the solvent, the lower will be the work required. Compressor process The solvent circuit in the extraction process using a compressor is schematically presented in the T–S diagram in Fig. 1.14. After the extraction process, which is performed under constant conditions (3), the solute is separated from the SCF by isenthalpic throttling into the subcritical region (3–4) and this is followed by evaporation (4–5–1). The dissolved substances precipitate and the solvent, which is in a gaseous state, is then compressed to the pressure of the extraction, shown in Fig. 1.14 as an idealised, isentropic compression (1–2). During compression, the temperature of the SCF increases, therefore the solvent is cooled to the temperature of extraction (2–3). This completes the cycle, which may be repeated. The operating modes in Figs 1.13 and 1.14 are presented as examples. Solvent cycles may vary for different gases and depend upon the mode of H6
P5
H5
P4 P3
Temperature (T)
H4 H3
Pc
SCF
2
H2
H1
P2 P1
3
CP
G 1/5
L
4
L–G
Entropy (S)
Fig. 1.14 Gas extraction process in T–S diagram: solvent circuit in the compressor mode.
© Woodhead Publishing Limited, 2010
24 Separation, extraction and concentration processes operation, e.g. the stages involved in separating the extract from the solvent can be done by pressure decrease, temperature increase, temperature decrease or a combination of temperature and pressure change. The mass and heat losses should also be considered when calculating the balance for practical processes. 1.3.2 Separation of solute in extraction processes using SCF After the solubilisation of SCF soluble substances in the extractor, the dissolved substances must be separated from the solvent–SCF mixture. Several separation processes may be used, but the preferred option is to use processes in which no additional substances are introduced into the system, which have varying solvent powers for supercritical fluid, and where only small changes in the conditions of the extraction process are made. Adding substances to separate solute from the SCF increases the costs of separation, whereas re-compression and cooling conditioning of the SCF to operating extraction conditions, increases the energy costs of the process. One of main advantages of using SCF for extraction and fractionation of substances, is that by a stepwise reduction of solvent power, fractionation of extracts can be obtained. Separation of solute by reduced solvent power Reduction of pressure and temperature increase The mostly widely used process for the separation of solute from solvent, is reduction of the solvent power. A condensed phase is formed and the gaseous phase is separated. The solvent power of supercritical solvents depends upon pressure and temperature which influence the density of the SCF. Generally, the solubility of a solute increases with increasing density and decreases with decreasing density. A decrease in density can be obtained by decreasing pressure and/or increasing temperature. In many industrial operations, the concentration of solute in a solvent is reduced by decreasing the pressure and thus the density, whereas an increase in temperature increases the vapour pressure of the solute. In a low pressure range where the density decreases dramatically with a temperature increase, the concentration of solute in SCF decreases with increasing temperature when the vapour pressures of the substances are relatively low. The separation of a solute by temperature increase is not an efficient method (for example for separation of essential oils from SCF) when the vapour pressure of a solute is high and increases with a rise in temperature. High pressure, where density decreases slightly with an increase in temperature, will lead to a higher concentration of solute with increasing temperature (at constant pressure) in the SCF. Based on these observations, it is evident that the increase in temperature, which decreases the density and solvent power, should provide an efficient method for the separation of a solute with low or moderate vapour pressure within the low-pressure range. On the contrary,
© Woodhead Publishing Limited, 2010
Principles of supercritical fluid extraction and applications 25 for some substances, at certain process conditions, the solubility decreases with increasing pressure. For such systems, separation procedures of solutes by pressure reduction or reduction of the solvent power is not possible. It is evident that the design of efficient processes for the separation of solute from SCF by the reduction of solvent power is possible only when phase equilibrium data are available. After establishing thermodynamic conditions for reduced solubility in a separator, a condensed phase is formed and this is separated from the solvent with lower density in one, or several separators in series (in the last separator from the gaseous phase). Usually, such a series of separators will operate at different pressures and/or temperatures to fractionate the single components of the extract. The solute–solvent system must remain in the separator for a sufficient length of time to approximately achieve the phase equilibrium. Separation by expansion Under certain process conditions, the solute-loaded solvent phase is expanded in the separator where the solvent is a subcritical liquid. In this instance, both the gaseous solvent and the liquid solvent containing the solute are present in the separator, a three-phase mixture is present if the solute is insoluble or less soluble in the liquid and gaseous phases. Usually, the liquid solvent evaporates and the solute is removed from the separator, whereas the gaseous phase is reused. This set of separation sequences is effective for the separation of a solute from solute-loaded SCF, but it is extremely expensive owing to the high energy input required. For efficient separation of the condensed phase from low-density solvent, auxiliary devices are used to increase the separation of the solid or liquid phase from the low-density solvent. These include demisters consisting of wire mesh packing, deflectors and filters, and cyclones. Separation of solute and solvent by a mass separating agent The separation of solute from a solution by a mass separating agent is possible by: ∑ absorption, ∑ adsorption, ∑ use of membranes, or ∑ adding a substance of low solvent power. For separation of a solute from a solvent by absorption, the solvent circuit may operate at an almost constant pressure. The absorbing liquid must dissolve the solute and should not absorb the solvent. Separation of solute from the SC solvent by adsorption can be a very efficient process. This process, like that of absorption, can operate with practically no pressure and/or temperature drop. Therefore, both separation processes have a major influence on the economy of whole extraction process. Membrane separation processes may be efficiently used for the separation
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26 Separation, extraction and concentration processes of solute from gas phase, owing to the difference in molecular mass. The pressure drop through the membrane is relatively low, therefore the solvent regeneration costs are low. The solvent power of a SC solvent may be reduced by adding a substance of low solvent power (e.g. adding nitrogen to SC CO2 for several substances). A similar effect may be obtained, when the entrainer is separated from the SC solution (if entrainer was applied in extraction process). For separating the solute, the entrainer has to be removed by adsorption or absorption. The advantage of such a separation process is its ability to operate at an almost constant pressure, thus keeping costs to a minimum.
1.4 Extraction of solids using SCF 1.4.1 Design of extraction plant The fundamentals of the design criteria for SCF extraction of solids in relation to process and equipment are reviewed in this section. For successful engineering and/or design of a SC extraction process, the following parameters should be defined: ∑ ∑
specific basic data, thermodynamic conditions for the operation of extraction and separation process, ∑ mass transfer data for the system, and ∑ energy consumption by means of T–S diagrams. The size of extractors for a certain capacity of plant are defined by specific basic data and thermodynamic conditions. Mass transfer data determines the time solids remain in contact with the SC solvent, so determining the solvent circulating system, including energy (heating and cooling) consumption, and therefore the capacity of the pump/compressor, the surface of heat exchangers and the piping system. Specific data and pretreatment of solids For the design of industrial plant, the following information is essential: ∑ raw material specification, ∑ final product specification, ∑ required plant size, ∑ plant location, sometimes taking into account the prevailing local conditions. The specification of the raw materials influences the quality of the extracts obtained and the overall economy of the process. If the raw material is contaminated, either by waste products or plant protection materials, a specified level of purity in the extract cannot usually be achieved. On the other hand, if the raw material contains a low concentration of substances to be extracted, the economy of the process will be questionable. © Woodhead Publishing Limited, 2010
Principles of supercritical fluid extraction and applications 27 The final product specification is usually based on customer needs, expectations and specifications and has a considerable influence on the costs of the extraction process. The process conditions in the extraction stage should be determined in order to obtain maximum yield and high selectivity for the required substance at minimal separation costs. However, a multistep separation is required for the fractionation of extracts. From a previous study (Gamse and Marr, 2001) it is known that the extractor volume influences the investment costs for an extraction unit (investment costs are a logarithmic function of extractor volume). Location will determine the mechanical construction of an extraction plant [e.g. location in Good Manufacturing Practice (GMP) area or in an earthquake area]. Climate conditions will also influence the design of heating and cooling devices and electrical drives. Pretreatment of solids Substances which should be extracted with SC fluids may be divided in two categories: ∑ ∑
raw materials where only unwanted substances should be removed and the geometry of the raw material is maintained during the process, and raw materials where pretreatment is allowed and extracted substances are the main product.
Examples of the first group of materials are coffee and tea decaffeination, the defatting of some seed plant materials, and the separation of pesticides from rice and ginseng. In the first process, high selectivity for the separation of unwanted compounds is necessary and therefore the process parameters or isolation of substances have to be very precisely selected. These processes are used for high-volume and relatively high-value market products. For the other group of materials, any pretreatment may typically be used to achieve a high yield at a low solvent to feed ratio with a low energy input. Therefore, the raw material should be ground before extraction to increase the bulk density. Some non-ground material could have a bulk density of 150–250 kg m–3, whereas the same material when ground would have a bulk density of 350–500 kg m–3. The ground feed material has an average particle size of 0.4 to 0.8 mm. Smaller particles have a higher specific area and therefore a higher mass flux could be expected. However, the linear solvent velocity through the ground raw material is decreased and consequently a greater drop in pressure is obtained. The linear solvent velocity through the raw material is 0.7 to 0.8 cm s–1. Finely ground material may also cause clogging of sintered plates on extraction vessels or filters and tends to cause channelling. If the bulk density of a ground substance is less than 250 kg m–3, the material should be formed into pellets. Bulk density has a strong influence on the economics of the process. If the bulk density in the extraction vessel is too low, less raw material is loaded and the yield per batch is reduced. Therefore, for a given product capacity, the volume of the extraction vessels should be greater. The moisture content
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28 Separation, extraction and concentration processes of raw plant materials is between 8 to 15 wt%. The amount of water in plant material influences the economics of the process. In plant material having a higher water content, the relatively high solubility of water in CO2 means that polar substances are extracted at relatively low pressure. If the water concentration in the raw plant material is too low, the cells may shrink and hinder the mass transfer of substances which are extracted through the cell walls. The ideal moisture content for extraction of substances from plant material should be optimised by laboratory-scale tests. It is known that for the decaffeination of coffee and tea, the water content of green coffee beans should range from 35 to 45 wt% (Zosel, 1981), whereas in the extraction of astaxanthin from algae, it should be as low as possible (water–oil emulsions are formed and the yield of astaxanthin would therefore be very low). In the SC extraction process for the isolation of colorants and antioxidants from raw plant materials, the content of oils, essential oils and waxes has a particularly large influence on the extraction yield. These substances contained in raw plant material may act as entrainers for the extraction of valuable compounds (such phenomena could be observed for extraction of carotenoids from ground paprika, where, owing to the low solubility of carotenoids in pure CO2 even at ultra high pressure (1500 bar), the carotenoids could not be separated quantitatively from plant material. Therefore, in some processes, oils, essential oils and waxes are used as entrainers for extraction of valuable compounds. Thermodynamic data The solubility of an extracted substance in a supercritical solvent is critical to the economics of the extraction process. The highest possible loading of SC solvent should be achieved in the extraction stages of the process, whereas in the separation stage, the solubility of solute in solvent should be as low as possible. In some instances, the solvent power for solutes should be gradually reduced so the fractionated separation of solute from the solvent may be achieved. Details of the thermodynamic properties of the solute in SC solvents are described in section 1.2. Mass transfer Mass transfer data are essential for determining the extraction time and the capacities of pumps, compressors, and heat exchangers. The data on mass transfer should be optimised experimentally with raw plant material having defined humidity, particle size, and particle size distribution at defined optimal extraction pressure and optimal extraction temperature (data obtained from thermodynamic investigations) for a defined SC solvent. Based on mass transfer experiments, the mass of solvent per mass of feed should be determined for the highest possible yield of the substance to be extracted. Figure 1.15 presents the typical extraction curve for the isolation of a substance from solids. The extraction curve for the isolation of substances from solid raw plant material is generally divided into a period of constant
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Principles of supercritical fluid extraction and applications 29
Yield (wt%)
Diffusion-controlled mass transfer
Solubility-controlled mass transfer
Mass of solvent/mass of feed (S/F)
Fig. 1.15 Typical extraction curve.
extraction rate and a period of falling extraction rate (Sovova, 2005). Solubility controlled mass transfer prevails in the initial period of the process, whereas at higher S/F (longer extraction time), the mass transfer is controlled by diffusion. The diffusion and hydrodynamics influence the mass transfer rates. Diffusion The substance to be extracted may be located in the cells of the raw material or adsorbed on the surface of a solid matrix. Therefore, the mass transfer depends on the location of the substance. If it is adsorbed on the surface, the mass transfer rates are high and, vice versa, when the solid particles of the substances diffuse through cell walls, the mass transfer rates will be low. In some instances, the substances to be extracted also form complexes which require release by a chemical reaction (usually hydrolysis with water). For substances that do not form a chemical complex, the diffusion may be influenced by: ∑ ∑
reduction of particle size (reduction of diffusion path), and destruction of cells (by swelling or cracking cells, by ultrasound, by milling procedure).
Hydrodynamics Particle properties such as particle size, particle shape and particle size distribution may cause channelling, which reduces the flow of the SC solvent through the material to be extracted. Some materials swell during the extraction process and may reduce the flow of the SC solvent through the extractor. For some substances, the direction of flow of the SC solvent is important. In industrial scale processes, the flow will usually be from bottom to top, whereas under some process conditions, the opposite direction of flow
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30 Separation, extraction and concentration processes will give higher yields. Mechanical mixing in industrial scale plants is not feasible, but some flow restrictors that avoid channelling in the extraction bed are technically possible. An extended description on modelling mass transfer may be found in Brunner (1994). Energy consumption Energy consumption for an extraction process using SC solvents may be determined from the TS diagram as described in section 1.3. 1.4.2 Applications for extraction of solids using SCF There are numerous applications for extraction of solids using supercritical fluids. Several overviews are available (Brunner, 1994; Catchpole et al., 2009; Diaz-Reinoso et al., 2007; Eltringham and Catchpole, 2007; Fang et al., 2007; Gardner, 1993; King and Srinivas, 2009; Lack and Seidlitz, 1993; Lack and Simandi, 2001; Li, 2007; Meireles, 2007; Mendes, 2007; Moyler, 1993; Mukhopadhyay, 2007; Reverchon and Marco, 2007; Stahl et al. 1987; Teelli, 2009; Temelli et al., 2007. The web pages of equipment producers (Natex, Nova Swiss, Sitec, Uhde HPT) give references to their plants. From this information, it may be seen that the highest capacity equipment is installed for coffee and tea decaffeination. The second largest application is for the extraction of hop compounds. Extraction of spices for the production of oleoresins and the extraction of bioactive compounds from plants are also very widely used applications of SC fluids. One of the latest applications is the extraction of oil from the degumming residue to obtain highly concentrated and very pure lecithin (plant designed, manufactured and erected by Uhde HPT).
1.5 Extraction of liquids using SCF There are fewer industrial units using SCF for the separation of components from liquid mixtures. Extraction from liquid mixtures by SCF is similar to liquid–liquid extraction, where compressed gas is used instead of an organic solvent. In liquid–SCF extraction processes, pressure plays an important role. When pressure and/or temperature is changed, the physicochemical properties of the SCF, such as density, viscosity, surface tension, and dielectric constant are also changed. Selective extraction of components or the fractionation of total extracts is possible by the use of different gases for the different processes and/or by changing the process parameters. Another advantage is that depending on the feed material, the density difference between the two counter-current flowing phases can be adjusted. One of the most important advantages of using supercritical fluids is the simplicity of solvent regeneration in comparison with liquid–liquid extraction,
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Principles of supercritical fluid extraction and applications 31 where in most cases, a re-extraction or distillation step is necessary; this step is energy consuming, thereby making the process more costly. Heat treatment of the extract or raffinate phase may cause the degradation of heat-sensitive substances. In extraction plants where SCFs are used, solvent regeneration is achieved by changing the pressure and/or the temperature after the extraction stage, thus changing the density and therefore the solvent power of the gas, which can be recycled after the separation of the solute. Compared with the extraction of solids with SCF, liquids may be continuously introduced and withdrawn from the high-pressure extraction unit. This offers the advantage of higher throughputs in continuous operating counter-current processes. There have been some laboratory-scale studies on extraction in systems using liquid/supercritical fluid. Data on binary systems liquid/SCF can be found, but there is less information on systems using liquid/liquid/supercritical fluid, which is necessary for the design of extraction processes of liquid mixtures with supercritical fluids. 1.5.1 Operation methods and apparatus As in conventional continuous liquid–liquid extraction in liquid/sub- or supercritical solvent, several operating modes of extraction are available. Single-stage extraction is the simplest and is used in systems where the separation factors for a solute are high. Multistage separation is necessary when the separation factor between the components is in the order of 1–10. Various modes of operation in multistage processes are used, such as multistage crossflow in which a relatively low loading of solvent with the extract is obtained in each stage. In multistage counter-current extraction, high loading of solvent with extracts is possible and a different configuration of the apparatus is possible. Counter-current liquid/sub- or supercritical fluid extraction (SFE) may be modelled by the use of typical commonly used basic equations: mass balance, energy balance, equilibrium distribution coefficients, and mass transfer rate equations. For extraction, the following data are necessary: ∑ ∑
determination of the number of theoretical stages/transfer units, size and type of a separation device in respect to separation performance, ∑ design of the solvent cycle. From the above facts and experimental data, the costs of separation using liquid/sub- or supercritical processes may be determined. The costs per tonne of the feed are influenced by the throughput and mode of operation (batch processes have higher operating costs, whereas in continuous processes the costs are lower) and are in the range from c. 607/kg feed at a throughput of c. 200 tonne per year in batch processes, down to approximately 0.067/ kg feed at a throughput of c. 60,000 tonne per year for continuous process. Brunner (2009) reported that the usual methods of calculation yield results
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32 Separation, extraction and concentration processes with an error of ±30%. Even after completion of a project, it is unlikely that the margin of error in determining costs will be less than 5%. 1.5.2 Applications The applications of liquid/sub- or supercritical fluid extraction are numerous and are used in the separation of ethanol from water (Hsu and Tan, 1994; Knez et al., 1994), the separation of aromas in various alcoholic beverages (Gamse et al., 1999), the separation of components from citrus oils (Knez, 1989) and for the purification of tocopherols (Fleck et al., 2000). The separation of caffeine from CO2 is used widely in decaffeination processes. In Fig. 1.16, the high-pressure column of an industrial-scale decaffeination process is shown. In the future, further limitations on the use of organic solvents and the demands of new applications will be the deciding factors for sustainable processing.
1.6 Conclusion SCF-based technologies offer important advantages over organic solvent technology, such as ecological friendliness and ease of product fractionation. The extraction of hop components and the decaffeination of tea and coffee are the largest scale extraction processes using sub- or supercritical solvents to be used industrially. There are also several industrial plants in operation
Fig. 1.16 Column for liquid–SCF extraction under preparation for pressure test (length 12 m, diameter 0.8 m, operation pressure 500 bar, operation temperature 100 °C. (Photo: courtesy of Uhde HPT, Hagen, Germany.)
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Principles of supercritical fluid extraction and applications 33 for the extraction of spices for the food industry and of natural substances for use in cosmetics. There are fewer industrial units engaged in the separation of components from liquid mixtures using sub- or supercritical fluids. The main advantages of using SCF for the isolation of natural products are: solvent-free products, no by-products and a low temperature in the separation process. In addition, the processes may be easily linked with direct micronisation and crystallisation from SC CO2 by fluid expansion. However, the most important advantage of the use of SCF, is the selective extraction of components or the fractionation of complete extracts. This is made possible by the use of various gases for the isolation/fractionation of components and/or by changing the process parameters. In addition to the wide use of gas for sub- or supercritical extraction, usually CO2, other sub- or supercritical solvents are also used. Sub- and supercritical CO2 and supercritical H2O are non-carcinogenic, non-toxic, non-mutagenic, nonflammable and thermodynamically stable. In addition, CO2 does not usually oxidise substrates and products, thus allowing the process to be operated at low temperatures. Water is currently the cheapest solvent and several substances are highly soluble in water. Further research on the use of subor supercritical water for the isolation and fractionation of substances is underway. One of the major process benefits is derived from the thermophysical properties of SCF: high diffusivity, low viscosity, density, and the dielectric constant of SCF, all of which may be fine tuned through changes of operating pressure and/or temperature. The limitation on further applications for obtaining extracts through highpressure technology is the price of the product; i.e. the price is relatively higher than for conventionally obtained extracts. The legal limitations on solvent residues and solvents (in products used for human applications) and the isolation/fractionation of special components from total extracts in combination with different formulations (controlled release for example) (Reverchon 2009; Weidner, 2009), chromatography (Taylor, 2009) and sterilisation processes, will lead to an increase in the use of dense gases for extraction applications. Figure 1.17 presents the number of compressed fluid extraction plants, by region (Lütge and Schuetz, 2007). It is evident that the number of extraction units will increase with time and the most ‘dense’ areas are in Europe and Asia, where a number of newly installed SCF extraction units may be observed. Figure 1.18 shows the pressures used for research and development. There is no clear trend towards higher or lower pressure. However, as part of research and development and in the filing of patents for SC extraction, both higher and lower pressures are used. Many substances have limited solubility at moderate pressure and the solubility of several substances increases by some orders of magnitude when using pressures over 500 bar. Because of this, the ultra high pressure range (up to 2500 bar) enables additional fractionation of substances in the extract
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34 Separation, extraction and concentration processes 50 Europe
America
Asia
Australia
45 40 35
Number
30 25 20 15 10 5 0 1980
1985
1990
Year
1995
2000
2005
Fig. 1.17 Number of SFE plants having a total extraction volume >500 L by region. (Source: Schuetz Consulting/Uhde HPT.) 800 700
Pressure (bar)
600 500 400 300 200 100 0 1980
1985
1990
1995 Year
2000
2005
2010
Fig. 1.18 Maximum pressure of SFE processes. (Source: Schuetz Consulting/Uhde HPT.)
by pressure- and/or temperature-dependent precipitation. In Fig. 1.19 an ultra high pressure extraction unit operating up to 2500 bar is presented. The evaluation of costs for several plant materials processed by ultra high pressure extraction using SC CO2 is shown in Fig. 1.20. It is clear that with © Woodhead Publishing Limited, 2010
Principles of supercritical fluid extraction and applications 35
E2
S1
E1
S2
Fig. 1.19 Ultra high pressure extraction unit (operating pressure 2500 bar). (E1 and E2, extractors; S1 and S2, separators). (Photo: courtesy of Uhde HPT, Hagen, Germany.) 2.50
Specific processing costs (7/kg)
Operating costs (7/kg) Investment (7/kg) 2.00
1.50
1.00
0.50
0.00
300
500 1000 Pressure (bar)
1500
Fig. 1.20 Specific processing costs versus operating pressure (Lütge et al., 2007)
an increased operating pressure in a SC extraction plant (retaining the same capacity of equipment), the total processing costs decrease. In the future, further limitations on the use of organic solvents, new applications of several substances, changing customer requirements, sustainable © Woodhead Publishing Limited, 2010
36 Separation, extraction and concentration processes production and processing of substances, will all lead to new developments in high-pressure processing. We can be sure that advances in the field of high-pressure research into cheap and environmental friendly solvents, such as CO2 and some other gases and sub- or supercritical water, will open up new pathways for substances and products produced at high pressure.
1.7 References Y. Adachi and H. Sugie, Fluid Phase Equilib. 28(2) (1986) 103–118. A. Anderko, Fluid Phase Equilib. 61 (1990) 145–225. A. Anderko, in: J.V. Sengers, R.F. Kayser, C.J. Peters, H.J. White (Eds.), Equations of state for fluids and fluid mixtures, Elsevier, Amsterdam, 2000, pp. 75–126. J.D.S. Arons and G.A.M. Diepen, Rec. Trav. Chim. Pays-Bas, 82 (1963) 249–256. A. Bertuco and G. Vetter, High pressure technology: fundamentals and application, Industrial Chemistry Library, volume 9, 2001. G. Brunner, Counter-current separations, J. Supercrit. Fluids 47 (2009) 574–582. G. Brunner, Gas extraction. An introduction to fundamentals of supercritical fluids and the application to separation processes, Darmstadt: Steinkopff; New York: Springer, 1994. O.J. Catchpole, S.J. Tallon, W.E. Eltringham, J.B. Grey, K.A. Fenton, E.M. Vagi, M.V. Vyssotski, A.N. MacKenzie, J. Ryan, Y. Zhu, The extraction and fractionation of specialty lipids using near critical fluids, J. Supercrit. Fluids 47 (2009) 591–597. R.P. Danner and P.A. Gupte, Fluid Phase Equilib. 29 (1986) 415–430. B. Diaz-Reinoso, A. Moure, H. Dominguez, and J.C. Parajó, Antioxidant extraction by supercritical fluids, in J.L. Martinez, Supercritical fluid extraction of nutraceuticals and bioactive compounds, CRC Press, 2007, 275–304. R. Dohrn, Berechnung von Phasengleichgewichten, Friedr. Vieweg & Sohn Verlagsgesellschaft mbH, Braunschweig, 1994. R. Eggers, Supercritical fluid technology in oil and lipid chemistry, AOCS Press, 1996, 35–62. W. Eltringham and O.J. Catchpole, Processing of fish oils by supercritical fluids, in J.L. Martinez, Supercritical fluid extraction of nutraceuticals and bioactive compounds, CRC Press, 2007, 141–188. T. Fang, M. Goto, M. Sasaki, and D. Yang, Extraction and purification of natural tocopherols by supercritical CO2, in J.L. Martinez, Supercritical fluid extraction of nutraceuticals and bioactive compounds, CRC Press, 2007, 103–140. U. Fleck, G. Brunner and R. Karge, Purification of synthetic crude tocophenol acetate by means of supercritical fluid extraction, in: Proceedings of the 5th International Symposium on Supercritical Fluids, Atlanta, April 2000. T. Gamse, I. Rogler and R. Marr, 1999, Supercritical CO2 extraction for utilisation of excess wine of poor quality, J. Supercrit. Fluids 14 (1999) 123–128. T. Gamse and R. Marr in A. Bertucco, G. Vetter. High pressure process technology: fundamentals and applications, Industrial Chemistry Library, volume 9, 2001, 383. D.S. Gardner, Commercial scale extraction of alpha acids and hop oils with compressed CO2, in M.B. King and T.R. Bott, Extraction of natural products using near-critical solvents, Chapman & Hall, 1993, 84–100. J.-H. Hsu and C.-S. Tan, Separation of ethanol/water solution with supercritical CO2 in the presence of a membrane, in S.S.H. Rizvi, Supercritical fluid processing of food and biomaterials, Chapman & Hall, 1994, 114–122. M.J. Huron and J. Vidal, Fluid Phase Equilib. 3 (1979) 255.
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Principles of supercritical fluid extraction and applications 37 P.G. Jessop and W. Leitner, Chemical synthesis using supercritical fluids, Wiley-VCH, July 1999. J.W. King and K. Srinivas, Multiple unit processing using sub- and supercritical fluids, J. Supercrit. Fluids 47 (2009) 598–610. Ž. Knez, Separation of components of citrus oils – industrial project, unpublished data, 1989. Ž. Knez, F. Posel and I. Krmelj, High pressure extraction of organics from water, in S.S.H. Rizvi, Supercritical fluid processing of food and biomaterials, Chapman & Hall, 1994, 181–186. Ž. Knez and M. Škerget, Phase equilibria of the vitamins D2, D3 and K3 in binary systems with CO2 and propane, J. Supercrit. Fluids 20 (2001) 131–144. P.H. van Konynenburg and R.L. Scott, Critical lines and phase equilibria in binary van der Waals mixtures, Phil. Trans. Roy. Soc. London, Ser. A 298 (1980) 495. E. Lack and H. Seidlitz, Commercial scale decaffeination of coffee and tea using supercritical CO2, in M.B. King and T.R. Bott, Extraction of natural products using near-critical solvents, Chapman & Hall, 1993, 101–139. E. Lack and B. Simandy in A. Bertucco and G. Vetter. High pressure technology: fundamentals and application, Industrial Chemistry Library, volume 9, 2001, 537–575. S. Li, Application of supercritical fluids in traditional chinese medicines and natural products, in J.L. Martinez, Supercritical fluid extraction of nutraceuticals and bioactive compounds, CRC Press, 2007, 215–242. T.W. de Loos, On the phase behaviour of asymmetric systems: The three-phase curve solid–liquid–gas. J. Supercrit. Fluids 39 (2006) 154–159. D. Luedecke and J.M. Prausnitz, Fluid Phase Equilib. 22(1) (1985) 1–19. C. Lütge and E. Schuetz, Market trends and technical developments in high pressure technology, in: Ž. Knez and M.J. Cocero, Proceedings of the 5th international symposium on high pressure process technology and chemical engineering, June 24–27, 2007, Segovia, Spain. C. Lütge, M. Bork, Ž. Knez, M. Knez Hrn�i�, M. Krainer, Ultra high pressure dense gas extraction and fractionation, in: Ž. Knez and M.J. Cocero, 5th International symposium on high pressure process technology and chemical engineering, June 24–27, 2007, Segovia, Spain. R. Marr and T. Gamse, High pressure technology: fundamentals and application, Industrial Chemistry Library, volume 9, 2000, 396–402. M.A. McHugh and V.J. Krukonis, Supercritical fluid extraction: principles and practice, Butterworths, Stoneham, 1986. M.A. Meireles, Extraction of bioactive compounds from Latin American Plants, in J.L. Martinez, Supercritical fluid extraction of nutraceuticals and bioactive compounds, CRC Press, 2007, 243–274. R. Mendes, Supercritical fluid extraction of active compounds from algae, in J.L. Martinez, Supercritical fluid extraction of nutraceuticals and bioactive compounds, CRC Press, 2007, 189–214. D.A. Moyler, Extraction of flavours and fragrances with compressed CO2, in M.B. King and T.R. Bott, Extraction of natural products using near-critical solvents, Chapman & Hall, 1993, 140–183. M. Mukhopadhyay, Processing of spices using supercritical fluids, in J.L. Martinez, Supercritical fluid extraction of nutraceuticals and bioactive compounds, CRC Press, 2007, 337–366. J.P. O’Connell, W. Weber and G. Brunner, Measurement and thermodynamics of triglyceride melting in near-critical fluids, in: Paper Presented at the AIChE Annual Meeting, Indianapolis, USA, 2003. A.Z. Panagiotopoulos, R.C. Reid, S. Watanasiri, Fluid Phase Equilib. 29 (1986) 525.
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38 Separation, extraction and concentration processes J.M. Prausnitz, R.N. Lichtenthaler and E. Gomes de Azevedo, Molecular thermodynamics of fluid phase equilibria. Prentice-Hall, Inc., Englewood Cliffs, New Jersey, 1986. E. Reverchon and I. De Marco, Essential oils extraction and fractionation using supercritical fluids, in J.L. Martinez, Supercritical fluid extraction of nutraceuticals and bioactive compounds, CRC Press, 2007, 305–336. E. Reverchon, R. Adami, S. Cardea, G. D. Porta, Supercritical fluids processing of polymers for pharmaceutical and medical applications, J. Supercrit. Fluids 47 (2009) 484–492. R.J. Sadus, High pressure phase behavior of multicomponent fluid mixtures, Elsevier Amsterdam, 1992. S.I. Sandler and H. Orbey, Mixing and combining rules; in: J.V. Sengers, R.F. Kayser, C.J. Peters, H.J. White (Eds), Equations of state for fluids and fluid mixtures, Elsevier, Amsterdam, 2000, pp. 321–357. R. Sandoval, G. Wilczek-Vera, J.H. Vera, Fluid Phase Equilib. 52 (1989) 119. J. Schwartzentruber and H. Renon, Ind. Eng. Chem. Res. 28 (1989a) 1049. J. Schwartzentruber, H. Renon, Fluid Phase Equilib. 52 (1989b) 127. H. Sovova, Mathematical model for supercritical fluid extraction of natural products and extraction curve evaluation, J. Supercrit. Fluids 33 (2005) 35–52. E. Stahl, K.-W. Quirin, D. Gerard, Verdichtete Gase zur Extraktion und Raffination, Springer-Verlag, 1987, 82–225. M Škerget, Z. Novak-Pintari�, Ž Knez and Z. Kravanja, Estimation of solid solubilities in supercritical carbon dioxide, Peng–Robinson adjustable binary parameters in the near critical region. Fluid Phase Equilib. 5086 (2002) 1–22. L.T. Taylor, Supercritical fluid chromatography for the 21st century, J. Supercrit. Fluids 47 (2009) 566–573. F. Temelli, M.D.A Saldaña, P.H.L. Moquin, and M. Sun, Supercritical fluid extraction of specialty oils, in J.L. Martinez, Supercritical fluid extraction of nutraceuticals and bioactive compounds, CRC Press, 2007, 51–102. F. Temelli, Perspectives on supercritical fluid processing of fats and oils, J. Supercrit. Fluids 47 (2009) 583–590. W.H. Tuminello, G.T. Dee and M.A. McHugh, Macromolecules 28 (1995) 1506. E. Weidner, V. Wiesmet, Z. Knez, M. Skerget, Phase equilibrium (solid–liquid–gas) in polyethylene glycol–carbon dioxide systems, J. Supercrit. Fluids 10 (1997) 139–147. E. Weidner, High pressure micronization for food applications, J. Supercrit. Fluids 47 (2009) 556–565. D.S.H. Wong and S.I. Sandler, AIChE J. 38 (1992) 671. K. Zosel, Process for the decaffeination of coffee, US Patent 4247570, January, 1981.
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Principles of pressurized fluid extraction and applications 39
2 Principles of pressurized fluid extraction and environmental, food and agricultural applications C. Turner and M. Waldebäck, Uppsala University, Sweden
Abstract: The main principles of pressurized fluid extraction (PFE), including basic extraction theory and the effects of solvent selection and temperature variations are discussed. Methods for achieving selectivity during the extraction are described, and future trends explored. Examples of applications include extraction methods developed for environmental and food analysis. Some advice on building equipment in the laboratory is given. Key words: food analysis, pressurized fluid extraction, PFE, environmental analysis, solvent extraction.
2.1 Introduction When developing new products, processes and technologies, it is important to strive for sustainable development. This is especially important because chemists, chemicals, and the chemical industry as a whole are commonly regarded to be the cause of many of the current environmental problems. To ensure sustainability, new processes and new techniques should be studied from a life cycle point of view. Life cycle assessment (LCA) is a methodology in which the entire life cycle of a product or utility effect is analyzed, such as extraction and processing of raw materials, production, distribution, use, consumption and disposal as well as the potential ecological effects. In addition, energy conversions occurring in a life cycle and the resulting burden on the environment are assessed. The indoor environment and health perspective is as important as the outdoor environment. Chemicals should be used in smaller amounts and the ones used should be less hazardous.
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40 Separation, extraction and concentration processes One process in which large volumes of organic solvents are used is the procedure of extraction. Traditional liquid–liquid extraction (LLE), also known as solvent extraction, and liquid–solid extraction (LSE) are techniques in which different compounds can be separated from each other based on their relative solubility. For a solid sample, the separation of a substance from the mixture occurs by dissolving that substance in an appropriate solvent. The extraction process usually requires several hours or even several days to perform, depending on the extraction temperature. These processes are slowly being replaced by more attractive alternatives. The most widely used extraction techniques today are still Soxhlet (developed in 1879) and sonication extraction (from 1960). These classical techniques are usually multi-step procedures based on exhaustive extractions from a sample matrix followed by successive clean-up steps before analysis. Such sample preparation procedures require large amounts of sample, sorbents and organic solvents, which are often hazardous and/or toxic, resulting in high costs of both purchase and disposal (Ramos et al., 2002). These methods also demand extensive manual handling, which often creates work-related health problems. During the past few decades new techniques have been developed and among the more successful ones are those employing high-diffusion liquids. Highdiffusion liquids are in this chapter defined as liquids of elevated temperature and pressure. By using high-diffusion liquids, the diffusion coefficient of the liquid is increased. This is the most effective way to increase the rate of the extraction process and decrease the required amount of organic solvent. Diffusion rates in liquids have been shown to increase about 2–20 fold upon increasing the temperature from 25 to 150 °C (Perry et al., 1984). Thus, the mass transfer rate can be increased and the whole extraction process becomes faster. Techniques that employ high-diffusion liquids with raised temperature and pressure are microwave-assisted extraction (MAE) and pressurized fluid extraction (PFE) (Majors, 1996). Another related extraction technique is supercritical fluid extraction (SFE), which usually employs carbon dioxide at pressures and temperatures above 74 atm and 31 °C, respectively, resulting in a liquid-like density in combination with high diffusion rates and low viscosity. In PFE, the solvent is kept in a liquid phase even at temperatures much above the atmospheric boiling point as a result of the applied pressure. This technique is also known as pressurized liquid extraction (PLE“), pressurized solvent extraction (PSE™), accelerated solvent extraction (ASE“) and enhanced solvent extraction (ESE). PFE was originally introduced at the Pittsburgh conference (Pittcon) in 1995 as ASE“ by the Dionex Corporation. The PFE technique utilizes the same basic principles as traditional liquid solvent extractions, but the extractions are carried out at higher temperatures and pressures. An increased temperature during the extraction gives more efficient extraction, and results in both time savings and lowered solvent consumption (Richter et al., 1996). The principle of PFE is similar to
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Principles of pressurized fluid extraction and applications 41 MAE, where the solvent is heated by microwave energy, although in the PFE technique both higher temperatures and pressures can be obtained, independently. In this chapter, the basic instrumentation of dynamic (continuous flow) and static (batch mode) PFE as well as its principles of operation are described. Furthermore, fundamental extraction theory is explained including illustrative examples and classical extraction models from literature. A guide to solvent selection is included, which on an elementary level guides the reader to an appropriate choice of solvent in various applications, depending on the chemistry of the sample matrix and analytes to be extracted. Effects of changing the extraction temperature, time and pressure in PFE are described, in addition to some remarks on how to obtain selectivity as well as accuracy and precision. One section is devoted to describing common applications for this technique. Finally, future trends for PFE in terms of sustainability, ‘like-nature’ applications, hyphenated techniques, downscaling and transfer to industry are depicted. The final paragraph also lists commercial vendors of PFE equipment and some advice for building equipment in the laboratory. The next subsection is largely based on a doctoral thesis written by one of the authors of this chapter (Waldeback, 2005).
2.2 Instrumentation and principles of pressurized fluid extraction 2.2.1 Basic instrumentation Extractions using a pressurized liquid above its atmospheric boiling point require open/close valves and/or pressure restrictors, in order to maintain pressure during the extraction. Two main types of instrumentation can be used: static PFE, which is a batch process with one or several extraction cycles with replacement of solvent in between; and dynamic PFE, in which the extraction solvent is continuously pumped through the extraction vessel containing the sample. A schematic of a static PFE instrument is shown in Fig. 2.1. Temperatures applied usually range from room temperature to 200 °C and pressures are generally between 35 and 200 bar. Filter paper is inserted into a stainlesssteel extraction cell followed by the sample, if necessary mixed with a drying agent. The cell is either loaded on a carousel and automatically placed in the oven, or, for simpler equipment, manually placed into the oven. Most equipment has a pump that fills the extraction cell with solvent. Commonly, a static valve in combination with a pressure relief valve control the pressure in the sample vessel during the static extraction; either by adding more solvent to the cell or by opening the static valve, whichever is needed to maintain the desired pressure. The first part of the extraction is a pre-heating step to reach thermal equilibrium. During this heating, thermal expansion of the
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42 Separation, extraction and concentration processes Purge valve
Pressure relief valve Extraction cell Oven
Pump
Solvent A Solvent B Static valve
Vent Solvent C Nitrogen
Waste vial
Solvent D Collection vial
Fig. 2.1 Schematic diagram of a PFE system (courtesy of Dionex Corporation).
solvent occurs and causes an increase in pressure within the cell. When the set values are achieved static extraction is performed during a selected time, typically 5–10 min. After the static time, part of the solvent in the extraction cell is replaced with fresh solvent, to start the next extraction cycle. After the last cycle, the sample cell is purged with an inert gas such as nitrogen to remove the remaining solvent from the cell and the lines to a vial that contains the extract. In some equipment, sequential extractions can easily be performed by repeating the procedure with a new solvent, and purging in the same or a different collection vial. Dynamic PFE is quite similar to static PFE, but requires a more sophisticated high-pressure or HPLC pump as well as a pressure restrictor rather than a static open/close valve. The equipment is similar to that used for ‘superheated water chromatography’ (Smith et al., 1999), but the column is replaced with an extraction cell and the tubing has slightly wider inner-diameter. Currently, there is no dynamic PFE equipment available on the market. In fact, there are only a few different types of commercial instruments available (Section 2.5.1), but many examples of home-made instrumentation have been described. Some of them perform both dynamic and static extractions at temperatures above 200 °C (Bautz et al., 1998; Hawthorne et al., 2000; Lou et al., 1997; Vandenburg et al., 1998). A short paragraph about building
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Principles of pressurized fluid extraction and applications 43 your own equipment can be found in Section 2.5.2, including Fig. 2.8 that shows a schematic of what a dynamic PFE system may look like. 2.2.2 Extraction strategy In order to perform efficient and smart extractions, an understanding of the matrix characteristics and the different steps of the extraction is of great importance. The nature of the sample matrix (water and organic/inorganic content) and its physical characteristics (homogeneity, porosity, particle size) should also be considered. To better understand the extraction process, two models showing the distribution of analytes in different types of sample matrices are illustrated below. Figure 2.2 is a conceptualization of an aggregate of matrix particles from a source-separated household waste, and the possible sites where analytes in this instance chlorinated paraffins, are expected to be found (Nilsson et al., 2001). Figure 2.2 works as a model showing the variety of possible (and likely) positions and status of analytes in many different types of sample matrices, i.e. the analyte can be: 1. 2. 3. 4. 5.
adsorbed at the surface of the matrix, dissolved in the pore solvent and/or adsorbed at the pore surface, dissolved/adsorbed in a micro/nano pore, chemically bonded to the matrix, or dissolved in the bulk solution.
A simpler extraction model can be described for extracting small organic compounds from polymer particles, for instance the antioxidant Irganox
5 1 2
4 3
Household waste particle Stagnant solvent layer
Fig. 2.2 Schematic of a household waste particle and some possible sites (for explanation of numbers see text) where the analyte (chlorinated paraffin) might be adsorbed or chemically bonded (Waldeback, 2005).
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44 Separation, extraction and concentration processes 1076 from the polymer LLDPE (Waldeback et al., 1998), Fig. 2.3. The extraction process of compounds from polymer particles, or other similar sample matrices, generally follows the following steps: 1. diffusion of the solvent into the matrix, 2a. desorption of the analytes from the matrix (including breaking of chemical bonds), 2b. solvation of the analyte into the extraction solvent, 3. diffusion of the analyte out from the matrix, and 4. diffusion of the analyte through the stagnant solvent layer and into the bulk solvent. Knowledge about distribution coefficients and distribution ratios are useful tools to provide guidance to the selection of solvent for the extractive separation process. Solubility of the analytes, their diffusivity in the solvent and matrix characteristics are the main factors to consider, when choosing a solvent for a successful extraction process. It is also important to understand the mass transfer mechanism across chemical/physical interfaces in order to design liquid/liquid and liquid/solid extraction processes. Toxicity and sustainability aspects of the solvent should also be considered. Additionally, in most of the analytical-scale applications, the concentration of the target molecule is very low, and thus the rate of the extraction is not limited by the analyte concentration in the extraction solvent, but rather determined by the rate of mass transfer out of the matrix. In order to perform a fast and complete extraction, a solvent has to be chosen that has the right chemical properties to dissolve and release the analyte, but should preferably not dissolve other solutes in the sample, i.e. the solvent power should not be higher than needed. Solubility theory has been discussed and proposed in classic works by J. Hildebrand, who combined the correlation between vaporization and intermolecular forces, van der Waal forces and hydrogen bonding, to the correlation between vaporization and solubility behaviour (Hildebrand and Scott, 1962; 1964). This model assumes that the same intermolecular attractive forces have to be overcome to vaporize a liquid as to dissolve an analyte. The term ‘solubility parameter’ d was described by Hildebrand as the square root of the cohesive energy density c, giving a numerical value indicating the analyte behaviour in a specific solvent, equation [2.1], where DH is the heat of vaporization (J mol–1) of the solvent, R the gas constant (J K–1 mol–1), T the temperature (K) and V is the molar volume of the analyte.
d = c = DH – RT [2.1] V Hansen (1967) took Hildebrand’s work further and assumed that the total cohesive energy is a linear addition of three components; dh (hydrogen bonding ability contribution), dd (dispersion coefficient contribution), and
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Principles of pressurized fluid extraction and applications 45 dp (polarity contribution). They are linked by equation [2.2], where dt is the total solubility parameter (Fitzpatrick and Dean, 2002).
d t2 = d h2 + d d2 + d p2
[2.2]
Fitzpatrick and Dean (2002) predicted a suitable solvent to extract persistent organic pollutants (POPs) from contaminated soil and certified reference material using the Hildebrand solubility parameter and confirmed the results by experiments. The ideal extraction solvent from the calculations was a mixture of acetonitrile and dichloromethane (1:1 v/v), perhaps the best solvent from the view of solvent power, but not from a health and environmental view. When choosing a solvent it is a good start to have the Hildebrand and Hansen’s theories in mind. However, the choice of solvent in a particular situation involves other factors apart from the solvent power. As described in Fig. 2.2 and 2.3, the solvent has to penetrate the matrix thoroughly, break the bonds between the matrix and the analytes of interest, help the dissolved analytes to diffuse out from the matrix and finally be dissolved in the extraction solvent. The process of solute transfer across an interface between two liquid phases may be rate-controlled by molecular diffusion, by motion of eddies, by irregular surface disturbances or even by chemical reactions in the bulk of a phase or in the interface region. Local velocities in the interface region could be of importance as well as other factors affecting the local conditions, not least the presence of surface-active agents. The dependence of matrix geometry has been reported for SFE of mineral oils applied to metal devices (Bjorklund et al., 1996). Irganox 1076
1
2 3
Polymer particle Stagnant solvent layer
4
Fig. 2.3 Schematic showing the extraction of Irganox 1076 from an LLDPE polymer particle (Waldeback, 2005).
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46 Separation, extraction and concentration processes Partitioning processes have a central role of concern in the extraction procedure. These involve partitioning of the analytes between the surface of the matrix and the solvent, as well as chemisorption of the analytes on active surface sites and within the solvent. Different matrices behave somewhat differently, e.g. polymeric samples usually build up a layer of stagnant liquid around the polymeric particles, as seen in Fig. 2.3, through which the analyte has to transfer into the extracting solvent. In this case, the partitioning of the analyte between the stagnant liquid and the extraction solvent has to be considered. Soils differ strongly in surface physicochemical properties and grain-size characteristics. Sediments, on the other hand, contain water having various types of bonding, from free water available for the plants, to water strongly bound to the particles, therefore a variety of equilibria take place. In milk powder, hard shells of lactose are formed during the short crystallization time when the milk is spraydried (Walstra and Jenness, 1984), and this may lead to difficult extraction of lipids and fat soluble vitamins. If the extraction solvent is nonpolar, it may be necessary to add a more polar alcohol such as methanol or ethanol to the sample before extraction (Turner and Mathiasson, 2000). There are only a few theoretical models suggested for the PFE technique. For instance, Vandenburg et al. applied the ‘hot ball’ model, originally described for SFE, for the extraction of additives from polymeric samples using PFE (Vandenburg et al., 1998). By plotting ln(m1/m0), where m1 is the mass of analyte remaining in the particle of radius r at time t, m0 is the initial amount of analyte and D the diffusion coefficient of the analyte in the solvent, a linear portion is given in equation [2.3] (Cotton et al., 1993). In this instance, when ln(m1/m0) is plotted against time, the line falls steeply initially and shortly after becomes linear where it follows equation [2.3] (Vandenburg et al., 1998):
ln(m1/m0) = – 0.4977 – (p2Dt/r2)
[2.3]
The physical explanation of the shape of the curve is that the analyte near the surface is rapidly extracted until a smooth falling concentration gradient is established across the particle. The extraction rate is then completely controlled by the rate at which the analyte diffuses to the surface. By plotting the amount of extracted analyte versus the extraction time for different solvents at different temperatures, the resulting curves showed a good fit to the ‘hot ball’ model. With a good kinetic model of the extraction process it would be possible to predict experimental parameters, and find out where the extraction is expected to be only diffusion dependent. However, the models mentioned here can at most be considered to give useful hints when developing new extraction methods based on PFE. There is no theoretical model that includes strong solute-matrix interactions caused by chemical bonding between the solute molecules and active sites on the matrix. In addition, sample matrices are seldom homogeneous, thus the penetration of the solvent is difficult to foresee. As a result, accurate descriptions are by
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Principles of pressurized fluid extraction and applications 47 no means easy to provide because of experimental difficulties, as well as the chemical and mathematical complexity of the total extraction process. Empirical approaches, in combination with multivariate chemometric methods, are the most widely used. 2.2.3 Solvent selection As mentioned in Section 2.2.2, a solvent has to have the right chemical properties to desorb and dissolve the analyte preferably without dissolving other solutes in the sample, i.e. the solvent power should not be higher than needed. Richter et al. (1996) have suggested the conventional LLE solvent to be used when developing a new PFE method. Generally, when choosing a solvent the rule-of-thumb is ‘like dissolves like’, i.e. polar solvents dissolve polar analytes, and nonpolar solvents dissolve nonpolar analytes. In addition, the dipole moment and/or dielectric constants of solvents are useful for selecting an appropriate extraction solvent, see Table 2.1. Other important aspects to consider when facing a new extraction problem/challenge is in how the analyte of interest is dissolved or attached Table 2.1 Chemical properties of common extraction solvents (from: Handbook of chemistry and physics, CRC press, 52nd Edition, 1971–1972; and Handbook of chemistry and physics, CRC Press, 60th Edition, 1979–1980) Solvent Boiling Density point (°C) (g mL–1)
Vapour Dipole pressure moment at 20 °C (debye) (mbar)
Dielectric constant at 20 °C (mbar)
Polar, protic solvents Water 100.0 Methanol 64.9 Ethanol 78.5 n-Propanol 97.1 2-Propanol 82.4 n-Butanol 117.7
23 128 44 – 43 –
1.85 1.70 1.69 1.68 1.66 1.66
80.2 32.2 24.3 20.1 18.3 17.8
3.96
47.2
1.000 0.791 0.789 0.803 0.785 0.810
Polar, aprotic solvents Dimethyl sulfoxide 189 (DMSO) Acetonitrile 81.6 Acetone 56.2 Dichloromethane 40.0 Tetrahydrofuran 66 Ethyl acetate 77.1
0.786 0.790 1.326 0.886 0.894
93 233 453 200 97
3.92 2.88 1.60 1.63 1.78
37.5 20.7 9.1 7.4 6.0
Nonpolar solvents Chloroform Diethyl ether Toluene Benzene n-Hexane
1.498 0.714 0.867 0.879 0.660
211 507 29 108 160
1.0 1.15 0.4 0 0
4.8 4.3 2.4 2.3 1.9
61.7 34.5 110.6 80.1 68.9
1.092
0.56
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48 Separation, extraction and concentration processes to the matrix, see Fig. 2.2 and 2.3, and also the type and strength of the intermolecular forces between the solvent molecules, which partly can be elucidated by the ΔHvap value, alternatively by the vapour pressure of the solvent (Table 2.1). When dissolving a solute, the intermolecular forces, i.e. hydrogen bonding, dipole–dipole and/or van der Waals interactions between the solvent molecules have to break and form new bindings/interactions with the solute. In instances when the solute is chemically bonded to the matrix, as in Fig. 2.2 (site 4), the extraction solvent has to overcome such bindings and then form new solute–solvent interactions, which altogether will lead to a more stable system of lower energy. Methanol and acetonitrile, two organic solvents commonly used in extraction and separation processes, have somewhat similar dielectric constants, 32 and 37, respectively, but differ strongly with regard to intermolecular forces as methanol forms hydrogen bonds, whereas acetonitrile relies on dipole–dipole interactions. Hence, acetonitrile, which is a less environmentally friendly solvent than methanol, should only be used in cases where such dipole–dipole interactions are more suitable. What are the partitioning processes or equilibrium reactions taking place during the extraction? For example, when extracting samples of high water content, such as wet sediments, vegetables or animal muscles, using organic solvents, there is always a risk that a nonpolar solvent will not be able to penetrate water-sealed pores that contain the analyte, and thus result in a lower extraction yield and unrepeatable results. The more knowledge about the chemistry of the target solutes, pertinent coextractable (unwanted) compounds and chemical and physical properties of the sample matrix and solute–matrix interactions, the easier it is to make a good selection of extraction solvent for instance from the ones listed in Table 2.1. One interesting aspect with PFE is the possibility of choosing a solvent, apart from very low or high pH solvents that may not be applicable in all kinds of equipment owing to possible corrosion of the tubings. In a collaborative study with an oil company, the objective was to predict the amount of process chemicals that could migrate into the sea if there was an oil discharge. By using a water solution of the same salinity as seawater, oil samples containing added process chemicals were extracted at different temperatures and times. The results could be used to predict what would probably happen in a real world situation (data not presented – only used internally at the company). This type of investigation is of course more environmentally relevant than a study performed with traditional extraction techniques based on organic solvents. How does the PFE technique and the solvent selection fit to a more sustainable chemistry? One of the advantages with PFE is that owing to the higher temperature and higher mass transfer during the extraction, more environmentally friendly solvents or solvent mixtures can be used according to Nilsson et al., 2001). As described above, the solvation power of a liquid increases with higher temperature and higher pressure. Furthermore, the
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Principles of pressurized fluid extraction and applications 49 dielectric constant decreases with increasing temperature, which implies that the solvent polarity can be tuned by changing the temperature. This is especially true for water, the dielectric constant of which decreases from e = 78 at 25 °C to e = 56 at 100 °C; e = 36 at 200 °C; and e = 8 at 400 °C, since hydrogen bonding between the water molecules become less and less substantial at higher temperatures (Hawthorne et al., 1994). This can be compared to other polar, protic solvents and nonpolar solvents in Table 2.1. Hence, water is a very interesting alternative as an extraction solvent, and can potentially replace many of the organic solvents conventionally used in solvent extraction. The first work describing the use of water at elevated temperature and pressure for extraction of nonpolar solutes is one by Hawthorne et al. (1994). In this work, polar organics (e.g., chlorinated phenols), low-polarity organics (e.g., PAH), and nonpolar organics (alkanes) were extracted from soils and sediments using water at temperatures ranging from 50 to 400 °C. In our own work, water at 120 °C and 50 bar was used as a solvent for the extraction of polyphenolic glycosides from onions, as compared with using a conventional water/methanol (1:1) mixture at 80 °C (Turner et al., 2006). The developed method was also faster, 15 min compared with 2 h for the conventional. In a continuous study, a life cycle assessment was conducted to assess environmental impacts of the two methodologies, showing that the production and use of methanol compared with water has a large impact in terms of carbon dioxide emissions and energy usage, whereas the difference in energy usage for heating the solvents in the two extraction methods was much smaller (Lindahl et al., 2010). If more nonpolar solutes are to be extracted, and these solutes are slightly thermolabile, then ethanol is an interesting alternative to water since not as high a temperature is needed to obtain the same solvent strength. For instance, betulin was extracted from birch bark using ethanol at 120 °C and 50 bar in only 10 min (Co et al., 2009). Initially, the study tested to see if water at elevated temperature close to 200 °C could be an appropriate solvent for betulin, because pressurized liquid water at this temperature has a dielectric constant of around 36, and this is similar to that of methanol at room temperature (e = 32, see Table 2.1), which is indeed a good solvent for betulin. However, it turned out that it is not sufficient to use the dielectric constant as a measure of solubility of a solute in a solvent, as other chemical properties are also important. For this reason, the total (sometimes called ‘relative’) solubility parameter taking dh, dd and dp into account (equation [2.2], Hansen, 2007) was plotted as a function of temperature for betulin, water and ethanol, showing that even at 250 °C, the solubility of betulin in water is estimated to be very low (Fig. 2.4), whereas ethanol at temperatures between 50 and 150 °C should be appropriate, which was confirmed by experiment (Co et al., 2009).
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50 Separation, extraction and concentration processes 50
Solubility parameter (MPa0.5)
45 40 35 30 25 20 15 Water
10
Betulin
5
Ethanol
0 0
50
100 150 Temperature (°C)
200
250
Fig. 2.4 Solubility parameters as a function of temperature for betulin, ethanol and water (Co et al., 2009).
2.2.4 Temperature effects An increase in temperature usually tends to promote solubility, as the thermal kinetic energy rises. Increasing the temperature also facilitates analyte diffusion and/or reduces interactions between analytes and the matrix by disrupting intermolecular forces such as van der Waal’s forces, hydrogen bonding and dipole attractions. Higher temperatures also decrease the viscosity of a liquid solvent, thus enabling better penetration of matrix particles. An increased temperature will also decrease the surface tension of the solvent, allowing the solvent to better ‘wet’ the sample matrix (Richter et al., 1996). Both lower viscosity and lower surface tension facilitate better contact of the solvent with the solutes and thereby enhance the extraction. Besides the selection of an extraction solvent, temperature can be considered as the second most important parameter in PFE. The advantages of using higher temperatures are described in every study published on PFE. As a general rule of thumb, a higher temperature gives a better extraction yield. It needs to be remembered that the range of solvents applicable in PFE is much wider than that of Soxhlet extraction. A poor solvent for Soxhlet extraction can, however, be a good solvent in PFE owing to the higher extraction temperatures used (Lou et al., 1997). Hence, as described below, by using a higher temperature, a more environmentally friendly and/or less toxic solvent can be used, for instance water or ethanol. However, other factors might hinder the use of the highest temperature even if it should give the best yield, e.g. as predicted in an experimental design. For example, Waldeback et al. (1998) found that the sample matrix (granules of LLDPE) started to dissolve/melt and block the tubings of the instrument depending on solvent and temperature, e.g. at temperatures above 75 °C © Woodhead Publishing Limited, 2010
Principles of pressurized fluid extraction and applications 51 using tetrahydrofuran as a solvent. Pihlstrom et al. (2002) found that, where pesticides from canola seed were extracted, the results from the screening study showed that the interpretation of the chromatograms was difficult owing to the large number of co-extracted compounds. It was obvious that higher temperatures gave more matrix peaks (Nemoto and Lehotay, 1998), which were most probably derived from co-extracted lipids. In another study (Waldeback et al., 2004), where squalene was extracted from olive biomass, the thermostability of the analyte was a limiting factor in the choice of extraction temperature. To determine the concentration of squalene in the olive biomass, an optimization study using experimental design was performed for the variables extraction temperature, extraction time and concentration of acetone in a mixture of acetone and 2-propanol. Significant factors were determined to be temperature, extraction time, the interaction between temperature and extraction time, and the square of the extraction time. This study showed that, at temperatures above 100 °C, the yield of squalene decreased with extraction times longer than 12–15 min, indicating that squalene either decomposed or reacted with other sample components or with the solvent. An example of extremely thermolabile compounds are anthocyanins, which are antioxidants found in vegetables and fruits, such as grapes, red cabbage, red onions, blueberries and red beets. PFE has been used to extract anthocyanins from red cabbage and red onions using water/ethanol/formic acid (94:5:1, vol%) at 99 °C and 50 bar (Arapitsas and Turner, 2008; Petersson et al., 2008, 2010). However, preliminary results show that anthocyanins extracted from red onion start to degrade after only a few minutes at a set temperature. Hence, extraction and degradation naturally occurs simultaneously in PFE. The kinetics of extraction and degradation depend on the nature of the sample matrix, on the chemical properties of the solutes to be extracted, and on temperature. Theoretical extraction curves were calculated using experimental extraction curves (with simultaneous degradation) combined with experimental degradation curves, as schematically described in Fig. 2.5 below. The conclusion was that, theoretically, a 20 to 30% higher yield of anthocyanins could be obtained from red onions if there were no degradation problems. This would require an extraction time of around 30 min (plus 8 min pre-heating time). In another study (Co et al., 2009), an increasing extraction temperature resulted in higher antioxidant activity in both water and ethanol extracts of birch bark. Furthermore, the antioxidant activities did not level out at high extraction temperature; hence, it appeared that some types of reactions occurred, and that these resulted in higher antioxidant activity. For instance, hydrolysis of polyphenolic glycosides leads to aglycons, which results in compounds of higher antioxidant activity because of the higher number of ‘reactive’ hydroxyl groups. Hence, an important conclusion is that a higher temperature may result in higher extraction yield of the target molecules, but also degradative reactions as well as extraction of more unwanted compounds, i.e. a less selective extraction.
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52 Separation, extraction and concentration processes
25
Peak area (a.u.)
20
15
10
5
0
0
20
40 60 Time (min)
80
100
Fig. 2.5 Calculation of an ideal extraction curve (continuous lines on top) from experimentally obtained data on degradation (dashed line) and extraction with degradation (continuous line with circles), respectively, for cyanidin-3-(6≤ malonoylglucoside) in red onion (Petersson et al., 2010).
Another type of degradative reaction that may occur in food and agricultural samples is caramelization of sugars. When water is used as a solvent at temperatures of around 160 °C and above, sugars such as glucose caramelize (Montilla et al., 2006). Depending on the length of the extraction time, this obstacle may cause erratic results as well as plugging of tubings and filters. If such high temperatures are to be used, it is advisable to carry out the extraction as fast as possible, to minimize the occurance of caramelization. Maillard reactions may also occur during PFE, involving reducing sugars and amino acids or proteins, but also flavonoids, ascorbic acid and other carbonyl compounds (Manzocco et al., 2000). Maillard polymers can be a major contributor to antioxidant capacity at higher temperature (>170 °C), when using for instance water or water/ethanol mixtures as extraction solvents in food applications (Howard and Pandjaitan, 2008). Hence, in food applications, to be on the safe side, it is advisable to use an extraction temperature of less than 160 °C. 2.2.5 Extraction time The third most important parameter after extraction solvent and temperature, and maybe the absolutely most interesting one, is extraction time. Several studies have shown that PFE offers faster extraction methods relative to those based on Soxhlet, sonication and SFE, and similar to MAE (Dean, 1996;
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Principles of pressurized fluid extraction and applications 53 Focant et al., 2004; Vandenburg et al., 1999). For instance, a comparison study was made for the extraction of PAHs from contaminated soil using Soxhlet, two different MAE methods, SFE and PFE (Saim et al., 1997). The Soxhlet method needed 24 h, the SFE method 35 min, the two MAE methods 20 min and the PFE method only 10 min, and Soxhlet extraction gave the highest recovery and the worst precision. In another study, Heemken et al. (1997) made a comparison of PFE and SFE with Soxhlet extraction of PAHs, aliphatic hydrocarbons and chlorinated hydrocarbons from marine samples. Statistical evaluations of accuracy and precision showed that equivalent results were achieved. The extraction time needed for the PFE method was 15 min, whereas SFE needed 90 min and Soxhlet 24 h. In general, during method development, extraction time and temperature are the most important parameters that need to be optimized once the solvent or solvent mixture has been chosen, as well as the mode of sample pre-treatment. Typically, extraction times are in the range of 10–15 min, plus a preheating step of 5–10 min depending on the temperature used and instrumental capabilities. However, if a dynamic extraction system is used (Section 2.5.2), the solvent can be preheated and continuously pumped through the sample without needing to preheat the cell. In commercially available PFE systems, dynamic extraction is mimicked using several short cycles of static extraction, replacing a certain volume of extraction solvent in between each cycle. The most recent Dionex PFE system, ASE-350®, can accomplish semi-dynamic extractions by step-wise replacement of small volumes of solvent continuously during the extraction. 2.2.6 Extraction pressure An increased pressure in the PFE technique is mainly applied to keep the solvents as liquids, at temperatures above their atmospheric boiling point. Most PFE studies observe no difference in extraction yield when the pressure is varied in the range of 34–204 bar (Lundstedt et al., 2000), although it has been reported (Richter et al., 1996) that a higher pressure rendered a higher extraction yield, when a standard mixture of PAH and polychlorinated hydrocarbons from a reference material was spiked onto wet silica of different pore sizes, using methylene chloride/acetone (1:1 v/v) at 100 °C as a solvent. Higher pressure gave higher extraction yield, when the pore size of the wet silica was 300 Å, but no difference in recovery was observed with the dry silica. It was suggested that higher pressure could probably facilitate the extraction of analytes trapped inside matrix pores because the pressure would force the solvent into the pores of the matrix that normally would not be contacted by solvents at atmospheric pressure. Pressurized flow would also aid in the solubilization of air bubbles so that the solvent could more easily penetrate the sample matrix. Kremer et al. (2004) also obtained higher yield of acidic herbicides in soil when pressure was increased from 69 to 138 bar, at 100 °C with dichloromethane as solvent.
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54 Separation, extraction and concentration processes 2.2.7 Selectivity during the extraction Selective extractions are in general obtained by changing the solvent or the solvent power. Under ambient conditions (room temperature and atmospheric pressure) the solvent power of a liquid is essentially constant regardless of extraction conditions. In PFE, selectivity can mainly be obtained by varying the nature of the solvent or the solvent mixture (Majors, 1999; Ramos et al., 2002). It has been pointed out that a potential disadvantage with PFE is that the extraction tends to be exhaustive, therefore leading to nonselective extractions (Reighard and Olesik, 1996). Hawthorne et al. (2000) compared the difference in color of the extracts, when soils contaminated with PAH were extracted by Soxhlet, SFE, PFE and pressurized hot water extraction (PHWE). The Soxhlet and the PFE extracts were much darker mainly owing to the greater amount of extracted compounds, i.e. the extractions were less selective. Hence, to avoid coextracting compounds with maintained high yield of the target analytes, it is essential to choose the most selective solvent for the analyte combined with an appropriate temperature, i.e. high enough for efficient extraction of the target compounds, but not too high to avoid coextraction of unwanted molecules. Extraction time is also an important factor: the minimum time necessary should be used to achieve targeted extraction rather than exhaustive extraction of other matrix solutes. However, by using adsorbents in the extraction cell as an in-line clean-up step, more selective PFE extractions can be obtained. For instance, activated acidic alumina has been employed to adsorb fat when PCBs were extracted from spiked freeze-dried fish tissue, and in this way better chromatogram separations were achieved (Ezzell et al., 1996). Sulfuric acid impregnated silica has similarly been used to adsorb the lipids from fat-containing food and feed when extracting PCBs to obtain lipid-free extracts (Sporring and Björklund, 2004). In our own study (Nilsson et al., 2001), when extracting chlorinated paraffins from household waste, clean extracts and chromatograms were obtained when the drying agents Hydromatrix and sodium sulfate were used, a result which is in agreement with other studies (Hawthorne et al., 2000; Ramos et al., 2002). On the other hand, the use of adsorbents such as XAD-2, XAD-4, and XAD-16 did not provide the removal of unidentified interfering compounds. In another work of ours, squalene and a-tocopherol have been selectively extracted from olive oil without coextraction of triglycerides when utilizing Amberlite XAD-16 as adsorbent (data not published). A liquid chromatogram of the olive oil before and after PFE is illustrated in Fig. 2.6. 2.2.8 Accuracy and precision Accuracy and precision obtained with the various PFE methods are similar to those accomplished with other extraction techniques such as Soxhlet, sonication, SFE and MAE (Helaleh et al., 2005; Saim et al., 1997). As with any extraction technique, accuracy and precision of the chemical analysis
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Principles of pressurized fluid extraction and applications 55
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Fig. 2.6 Liquid chromatograms of (a) olive oil diluted with ethanol/acetone and (b) a PFE extract of olive oil, obtained using XAD-16 as adsorbent and methanol as extraction solvent.
method heavily depend on how well the sample is homogenized, and if the collected sub-sample is an accurate representation of the entire sample under study. These problems with sample preparation can of course be minimized by taking out a larger number of sub-samples, i.e. replicates, for analysis. Another important variable is the particle size, i.e. too large a particle size leads to inefficient extraction and inaccurate results. This was demonstrated by Björklund et al. (1999) for sediment samples of different particle size distributions. Similarly, Isaac et al. (2005) in a study of the extraction of lipids from homogenized cod tissue obtained a higher yield compared with extracting intact wet cod muscle. The conclusion was that the proteins in the sample denaturated to a hard pellet during the extraction making the diffusion of the solvent into the matrix not efficient enough to extract all the lipids. On the other hand, too small a particle size led to a relatively compact sample, which in turn may result in channeling inside the sample and again inefficient extraction.
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56 Separation, extraction and concentration processes
2.3 Applications of pressurized fluid extraction Most of the studies on PFE deal with either environmental analysis or food and agricultural samples, targeting either pollutants or health-beneficial compounds such as antioxidants. Figure 2.7 shows the publication record from 1995 to 2008, taken from Web of Science. There is not enough space in this chapter to discuss in detail the wide range of different applications for PFE, hence, the reader is referred to a number of reviews of the field: on applications for environmental analysis (Giergielewicz-Mozajska et al., 2001; Nieto et al., 2008); food and drugs (Beyer and Biziuk, 2008; Carabias-Martinez et al., 2005; Herrero et al., 2006; Mendiola et al., 2007); and medicinal plants, herbs and agricultural (Huie, 2002; Wang and Weller, 2006). 2.3.1 Environmental applications Pesticides were extracted from fruits and vegetables (Adou et al., 2001; Barriada-Pereira et al., 2007; Blasco et al., 2005; Herrera et al., 2002; Pihlstrom et al., 2002; Tanaka et al., 2007; Wennrich et al., 2001) and from agricultural crops and soils (Hildebrandt et al., 2007; Otake et al., 2008; Popp et al., 1997; Schreck et al., 2008) using solvents and solvent mixtures such as acetone, acetone/dichloromethane, acetone/hexane, methanol, hexane saturated with acetonitrile, and acetonitrile and water, at temperatures between 60 and 130 °C. For polyhalogenated persistant organic pollutants such as polychlorinated biphenyls (PCBs), polybrominated diphenyl ethers (PBDEs), polychlorinated dibenzodioxins (PCDDs) and polychlorinated dibenzofurans (PCDFs), it is more common to use hexane, toluene/acetone and hexane/ acetone, at temperatures between 100 and 150 °C (Antunes et al., 2008; Bjorklund et al., 1999; Brandl et al., 2006), but water has also been used at temperatures between 250 and 300 °C (Yang et al., 1995). Number of publications
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Fig. 2.7 Publication record (Web of Science 2009-08-02, search term: ‘pressurized liquid extraction’ OR ‘accelerated solvent extraction’ OR ‘pressurized solvent extraction’ OR ‘subcritical water extraction’ OR ‘pressurized fluid extraction’). In total 1469 publications between 1995 and 2008.
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Principles of pressurized fluid extraction and applications 57 Organometallic compounds containing tin, arsenic and mercury have been extracted from sediments and soils (Beichert et al., 2000; Chiron et al., 2000), vegetables (Marcic et al., 2005) and seafood (Mato-Fernandez et al., 2007; Wahlen and Catterick, 2004) using methanol, water, methanol/water and methanol/ethyl acetate mixtures, at temperatures between ambient and 160 °C. There are also a few studies on the extraction of explosives from soil samples (Ragnvaldsson et al., 2007) and biological samples (Pan et al., 2005); and antibiotics from soil (Jacobsen et al., 2004; Schlusener et al., 2003) and meat samples (Berrada et al., 2008). Pharmaceutical applications are less common than environmental ones, and most of them consider the extraction of pharmaceutical drugs and their metabolites in environmental samples such as sediments, sludge and fish/ shell fish. For instance, pharmaceutical residues have been extracted from sludge using methanol/water (1:1) at 60 °C as solvent (Barron et al., 2008). There is also an increasing number of studies on endocrine disruptors in the environment, for instance, 17-b-estradiol has been extracted from soil using acetone/hexane (1:1) at 100 °C (Chun et al., 2005); and alkylphenolic compounds have been extracted from river sediment using methanol/acetone (1:1) at 50 °C as a solvent (Petrovic et al., 2002). 2.3.2 Food and agricultural applications Food and agricultural applications considering endogeneous compounds commonly involve the use of more environmentally friendly solvents such as water and ethanol, because these applications often aim at isolating healthbeneficial compounds for later use as food additives, or alternatively, the compounds that are extracted are fairly polar, hence stronger organic solvents are not needed. Exceptions are in applications concerning the extraction of lipids such as acylglycerols, sterols, terpenoids and essential oils, for which stronger organic solvents are generally employed, e.g. isopropanol, ethyl acetate and hexane. Flavonoids such as quercetin, kaempferol, catechin and anthocyanidins have been extracted from vegetables and herbal plants using pressurized hot water, ethanol or water/ethanol mixtures as solvents at temperatures between 50 and 160 °C (Arapitsas and Turner, 2008; Howard and Pandjaitan, 2008; Ibanez et al., 2003; Ollanketo et al., 2002; Turner et al., 2006). For instance, water and a 70:30 mixture of ethanol and water at temperatures of 50 to 190 °C were used to extract flavonoids from dried spinach (Howard and Pandjaitan, 2008). The results showed that the total phenolic content as well as the antioxidant capacity were the highest in extracts obtained with the highest extraction temperatures (170–190 °C) for both water and the water/ethanol mixture, although the solvent mixture was somewhat more efficient than pure water in extracting flavonoids. However, when separating the extracts into low (<1000 Da) and high (>1000 Da) molecular weight fractions, it turned out that at temperatures above 130 °C for water and
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58 Separation, extraction and concentration processes above 150 °C for ethanolic solvents, polymeric Maillard reaction products were responsible for the increase in antioxidant capacity. Howard and Pandjaitan (2008) concluded that, in order to extract flavonoids effectively, water should be used at temperatures between 50–130 °C and for ethanolic solvents, 50–150 °C. In one of our own studies, quercetin and its glycosides as well as isorhamnetin and kaempferol were extracted from yellow and red onion using water at 120 °C and 50 bar, employing three consecutive 5-min extraction cycles (Turner et al., 2006). Thermostable b-glucosidase was used to convert the polyphenolic glucosides into their respective aglycons within only 10 min of reaction, using water at 90 °C and pH 5 as a reaction media. Results obtained compared well with a conventional extraction method employing a 1:1 mixture of methanol and 2.4 M HCl at 80 °C and 2 h combined extraction/ reaction, followed by filtration. Phenolic compounds such as rosmarinic and carnosic acids, carnosol and methyl carnosate have been extracted by PHWE, ultrasonication-assisted methanol extraction, hydrodistillation, and maceration with 70% ethanol from sage (Ollanketo et al., 2002). It turned out that PHWE, performed at 100 °C, gave the highest antioxidant activity of the extracts. Furthermore, PHWE was faster than the conventional techniques used in this work. In another study (Ibanez et al., 2003), antioxidative compounds such as carnosol, rosmanol, carnosic acid, methyl carnosate, cirsimaritin and genkwanin have been extracted from rosemary leaves using subcritical water at temperatures between 25 and 200 °C, showing that the selectivity of the extraction can be tuned by varying the temperature. At 25 °C a high concentration of rosmanol was obtained in the extract, while at 200 °C, the dominant compound was carnosic acid. The antioxidant activity was high in all of the extracts obtained at 100, 150 and 200 °C, respectively (Ibanez et al., 2003). Phenolic acids have been extracted from fruits and vegetables using water, methanol and water/methanol mixtures, at temperatures ranging between 20 and 100 °C (Alonso-Salces et al., 2001; Chen et al., 2007a; Mukhopadhyay et al., 2006; Waksmundzka-Hainos et al., 2007). In general, the main concern for the extraction of phenolic acids as well as many of the flavanoids is the thermal stability, or rather the tendency to degrade during the extraction, as discussed in Section 2.2.4. Terpenoids including limonene, pinene, artemisinin, retinol, taxol, squalene, lycopene and carotene, have only limited water solubility even at elevated temperatures, hence these are extracted with, for instance, 2-propanol or methanol/ethyl acetate/light petroleum (1:1:1) at temperatures ranging from 40 to 190 °C (Breithaupt, 2004; Fojtova et al., 2008; Schaneberg and Khan, 2002; Waldeback et al., 2004). There are, however, also studies in which pressurized hot water is used as a solvent; Yang et al. (2007) used water at temperatures between 100 and 250 °C to extract a-pinene, limonene, camphor, citronellol, and carvacrol terpenoids from oregano and basil leaves. However, it was shown that thermal degradation occurred at all temperatures tested,
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Principles of pressurized fluid extraction and applications 59 although least severe at 100 °C – between 10 and 25% for a 30-min treatment in static mode. In a work by Breithaupt (2004), carotenoids were extracted by PFE and compared with a conventional solvent extraction, in both cases using a solvent mixture of methanol/ethyl acetate/light petroleum (1:1:1). The PFE extraction conditions were optimized by varying the temperature (25–80 °C) and the pressure (70–140 bar), giving the highest yield at 40 °C and 70 bar. Both methods were used for the extraction of a large number of different carotenoids from beverages, pudding mixes, cereals, cookies and sausages (Breithaupt, 2004). Acylglycerols and sterols have been extracted from eggs and egg-containing food products, using chloroform/methanol 2:1 (v/v) and hexane/isopropanol 3:2 (v/v), at various extraction temperatures and pressures (60 °C and 150 bar, 100 °C at 150 bar, and 120 °C and 200 bar (Boselli et al., 2001). Results showed that the hexane/isopropanol mixture employed at 60 °C and 150 bar could successfully be used to extract lipids from egg products. In another study, lipids were extracted from freshly ground corn kernels and ground rolled oats, using four different organic solvents: hexane, dichloromethane, isopropanol, and ethanol, at two temperatures, 40 and 100 °C (Moreau et al., 2003). The results showed that oat oil with the highest levels of digalactosyldiacylglycerol (DGDG) was obtained using ethanol at 100 °C. Comparison of the extraction effects with seeds of two different species of grains indicated that the extractability is greatly dependent on the polarity and temperature of the solvent. Tocopherols and tocotrienols are fat-soluble compounds extractable using solvents such as methanol and acetone. For instance, tocopherols and tocotrienols have been extracted from cereals using methanol as extraction solvent at 50 °C and 110 bar, using only one 5-min extraction cycle (Bustamante-Rangel et al., 2007). Other food and agricultural applications include for instance alkaloids (Mroczek and Mazurek, 2009); saponins (Chen et al., 2007b); coumarins (Waksmundzka-Hajnos et al., 2004); oligosaccharides (Bansleben et al., 2008); lignans (Smeds et al., 2007); and various essential oils (JimenezCarmona et al., 1999; Ozel et al., 2003; Schaneberg and Khan, 2002).
2.4 Future trends The authors believe that some of the most important future trends for PFE are: (i) switching to more environmentally friendly solvents; (ii) hyphenation and automation; (iii) miniaturization; (iv) mimic nature; and (v) transfer to industry.
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60 Separation, extraction and concentration processes All these trends are somewhat connected with the growing push for a more sustainable development. In order to minimize the use of organic solvents that need to be produced and then disposed of, both switching to more environmentally friendly solvents and miniaturizing the extraction equipment is vital. The use of solvents at elevated temperatures implies that more efficient extractions can be obtained, even when using more environmentally friendly solvents such as water and ethanol rather than the classical ones such as hexane and ethyl acetate. Hence, future trends will most probably involve switching from conventional organic solvent based extraction methods to PFE using water, ethanol and other ‘green’ alternatives. This has already been demonstrated (Burkhardt et al., 2005; Co et al., 2009; Crescenzi et al., 1999; Curren and King, 2001; Fernandez-Perez and de Castro, 2000; Hawthorne et al., 1998; Hyotylainen et al., 2000; Morales-Munoz et al., 2002; Smith, 2002; Turner et al., 2006; Yang and Li, 1999), but more applications are expected in the near future. Regarding miniaturization of PFE, we found only one example (Ramos et al., 2000), in which PAHs were extracted from soil and sediment using an extraction vessel constructed by a 10 mm ¥ 3 mm inner diameter stainlesssteel tubing. The only extraction solvent used was 100 mL of toluene at 200 °C and 150 bar, in an extraction process taking 10 min, with subsequent direct injection into large-volume injection gas chromatography–mass spectrometry (GC–MS). Future trends will probably include more examples of miniaturized PFE systems, probably also combined with analytical separation and detection systems. It has been recognized that in order to minimize the production of waste and emissions of carbon dioxide, it is advantageous to use as few steps in a process as possible, i.e. combined (hyphenated) methods and automation are relevant. There are several good examples showing that PFE can be coupled on-line with separation by HPLC (Hyotylainen et al., 2000; Li et al., 2000), but also with GC using a miniaturized PFE system (Ramos et al., 2000). Smith showed that subcritical water extraction can be coupled on-line with chromatography using subcritical water as the only mobile phase (Smith, 2002; Tajuddin and Smith, 2002). In the future, technology development enabling combination with several other separation and detection techniques is expected, as well as in situ analysis instrumentation. In order to better understand processes occurring in real life or in the environment, laboratory-scale methods that mimic nature in an efficient way are important. Thus, by using PFE, an interesting possibility is to simulate what could happen in a real life situation (in the human body, in the environment or within an industrial process). This can be accomplished using a suitable solvent, similar to the surroundings in question, for example water with appropriate pH and ion concentrations (such as seawater). Anaerobic conditions can be simulated using a nitrogen purge between sequential extractions. Hence, it is possible to predict what would happen if a specific chemical is used by a specific company in a specific environment
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Principles of pressurized fluid extraction and applications 61 and, in this way, obtain detailed information for an environmental regulatory impact analysis. Another challenge would be to follow the migration of any hazardous compound and their degradation products in soil, sediment, sludge or biota. Ongoing research in this area is expected to address important environmental issues. A final note on future trends of PFE is that several of the smaller-scale processes will be transferred to industry to replace conventional solvent extraction methods in chemical analysis applications, as well as scaled up for industrial processes in applications regarding isolation of interesting (valuable) compounds from the plant kingdom. For instance, a standard method (US EPA SW-846 Method 3545) for the extraction of PCBs from sediments using hexane/acetone (1:1) as a solvent at 150 °C has replaced a traditional Soxhlet method. Both new applications and replacement of technology in old applications are expected to be developed in the near future.
2.5 Sources of further information and advice 2.5.1 Commercially available equipment There is equipment available on the market that performs static PFE, with full or semi-automation. Currently, the following vendors have been identified: Dionex (www.dionex.com) products include ASE® 150 (one extraction vessel at the time, of sizes 1, 5, 10, 22, 34, 66 or 100 mL) and ASE® 350 (carousel with up to 24 extraction vessels of the same sizes as for the 150-system). A solvent controller allows mixing and delivery of up to three solvents. Temperature range applicable is 40–200 °C and the pressure is fixed to 1500 psi (103 bar). pH hardened pathway with Dionium™ components and extraction vessels make the instrument compatible with acid or alkaline pretreated sample matrices. ∑ Applied Separations (www.appliedseparations.com) has two extraction systems, one PSE™ and fast-PSE™, which operate with one and six in-parallel extraction vessels, respectively. Available vessel sizes are 11, 22 and 33 mL. An automated solvent dispenser is optional, and up to four different solvents can be handled. Temperature range is 50– 200 °C and maximum pressure is 150 bar. ∑ Fluid Management Systems (www.fmsenvironmental.com) market PFE™ systems that handle one to six samples (six modules) at the time in parallel, with the option of in-line clean up by column chromatography. Temperature and pressure ranges obtainable are 70–200 °C and 1500– 3000 psi (103–207 bars), respectively. Extraction vessels come in sizes ranging between 5–250 mL. ∑
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62 Separation, extraction and concentration processes 2.5.2 Designing your own equipment In our laboratory, we usually build our own equipment, since the ones commercially available are all based on static PFE and, in addition, operate within a strictly limited temperature range. Hence, in order to run dynamic PFE, even at temperatures above 200 °C, it is necessary to custom build PFE instrumentation. What we recommend is to use the following parts: (a) an optional heater plate set to a temperature of around 20–30 °C below the boiling point of the solvent; (b) a bottle containing the extraction solvent or solvent mixture; (c) an ordinary HPLC pump; (d) a long stainless-steel tubing that enables the solvent to reach desired temperature before passing through the extraction vessel; (e) a type K thermocouple temperature probe for measuring the temperature of the solvent just before entering the extraction vessel; (f) a stainless-steel extraction vessel (for instance a preparative HPLC column) in which the sample is placed before extraction; (g) a simple GC oven (only temperature control is needed, hence no need for injector or detector); (h) a cooling bath to bring down the temperature of the extractant to below the boiling point of the solvent; (i) a needle valve or a back pressure regulator; (j) collection vial(s); and (k) optional nitrogen gas to flush out the entire system after extraction. It is advisable to install an in-line filter to protect the needle valve/back pressure regulator. It is also recommended to install a pressure gauge for control of the pressure as well as a burst disc in case the pressure rises too high in the system. There are burst discs available for different pressures and solvents. Figure 2.8 shows a schematic of a home built dynamic PFE system. (c)
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Fig. 2.8 Schematic of a home-built dynamic PFE system. (a) Heater plate; (b) extraction solvent; (c) HPLC pump; (d) stainless steel tubing; (e) temperature probe; (f) extraction vessel; (g) oven; (h) cooling bath; (i) needle valve; (j) collection vial; (k) nitrogen gas.
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Principles of pressurized fluid extraction and applications 63 A factor to consider is which solvents and temperatures are allowable by the different parts of the extraction system. For instance, an HPLC pump cannot take too high a temperature, and especially not with some of the organic solvents. If water is used as a solvent, caution must be taken if temperatures above ca. 200 °C are used, since water becomes more and more corrosive at higher temperature. Hence, higher-quality nickel steel alloy, such as Hastelloy® that exhibits high resistance to corrosion, should be used. However, owing to obstacles such as caramelization and Maillard reactions occurring in food and agricultural samples, temperatures applied are rarely above 160 °C. Finally, if an organic solvent is used, safety arrangements around your homebuilt device, such as ventilation and burst discs connected to stainless-steel tubings leading to a secure waste container, should be provided.
2.6 Conclusions In industrial and governmental analytical and research laboratories, PFE is a promising extraction technique that easily replaces conventional techniques such as Soxhlet extraction – especially since the same solvents can be used. Compared with conventional extraction techniques, PFE is faster; more automated; and enables a ‘friendlier’ work environment. Furthermore, new PFE methods are fairly easy to develop by optimizing temperature and extraction time. Since PFE operates at elevated temperatures, it is possible to use less harmful solvents than in Soxhlet with maintained extraction efficiency. However, caution has to be taken regarding degradation of thermolabile target compounds during the extraction, even though the problems are usually minor if the extraction time and temperature are carefully optimized. Finally, selectivity can be obtained in PFE by firstly selecting an appropriate extraction solvent, secondly by varying the temperature of the solvent, which is especially effective for water, or thirdly by using adsorbents in the extraction cell. How does PFE fit into a sustainable development? PFE uses smaller amounts of organic solvent, thereby preventing unnecessary waste, because less organic solvent needs to be manufactured and less waste is thereby produced by the manufacturer. In addition, less organic solvent waste is produced at the laboratory and, in total, there will be a decrease in volatile organic carbon emissions and a reduced risk for photochemical smog formation. Accidental emissions to municipal wastewater plants will do less harm. By being able to substitute hazardous solvents with more environmentally friendly solvents such as ethanol and water, the toxicity is reduced with positive effects on both the environment and health in general. Creating high diffusion liquids requires a significant amount of energy, which is a disadvantage. However, because the extraction times are shorter, minutes compared with hours in Soxhlet and modified Soxhlet techniques, and, as these traditional solvent extraction techniques are usually performed © Woodhead Publishing Limited, 2010
64 Separation, extraction and concentration processes just below the boiling point, a significant amount of energy is saved. Lifecycle assessment should be conducted to calculate the overall energy usage in different methodologies.
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70 Separation, extraction and concentration processes Waksmundzka-Hainos, M., Oniszczuk, A., Szewczyk, K. and Wianowska, D. (2007), ‘Effect of sample-preparation methods on the HPLC quantitation of some phenolic acids in plant materials’, Acta Chromatogr, 19, 227–37. Waksmundzka-Hajnos, M., Petruczynik, A., Dragan, A., Wianowska, D. and Dawidowicz, A. L. (2004), ‘Effect of extraction method on the yield of furanocoumarins from fruits of Archangelica officinalis hoffm’, Phytochem. Anal., 15, 313–19. Waldeback, M., Jansson, C., Senorans, F. J. and Markides, K. E. (1998), ‘Accelerated solvent extraction of the antioxidant Irganox 1076 in linear low density polyethylene (LLDPE) granules before and after g-irradiation’, Analyst, 123, 1205–07. Waldeback, M., Senorans, F. J., Fridstrom, A. and Markides, K. E. (2004). Pressurized fluid extraction of squalene from olive biomass, in Modern extraction techniques for food and agricultural samples. C. Turner. Anaheim, CA, ACS Press. 926, 96–106. Waldeback, M. (2005). Pressurized fluid extraction: a sustainable technique with added values. PhD thesis, Uppsala University, Department of Physical and Analytical Chemistry, Uppsala, Sweden. Walstra, P. and Jenness, R., 1984, Dairy chemistry and physics, New York, John Wiley & Sons. Wang, L. J. and Weller, C. L. (2006), ‘Recent advances in extraction of nutraceuticals from plants’, Trends Food Sci. Technol., 17, 300–12. Wennrich, L., Popp, B. and Breuste, J. (2001), ‘Determination of organochlorine pesticides and chlorobenzenes in fruit and vegetables using subcritical water extraction combined with sorptive enrichment and CGC–MS’, Chromatographia, 53, S380–S86. Yang, Y., Bowadt, S., Hawthorne, S. B. and Miller, D. J. (1995), ‘Subcritical water extraction of polychlorinated biphenyls from soil and sediment’, Anal. Chem., 67, 4571–76. Yang, Y. and Li, B. (1999), ‘Subcritical water extraction coupled to high-performance liquid chromatography’, Anal. Chem., 71, 1491–95. Yang, Y., Kayan, B., Bozer, N., Pate, B., Baker, C. and Gizir, A. M. (2007), ‘Terpene degradation and extraction from basil and oregano leaves using subcritical water’, J. Chromatogr. A, 1152, 262–67.
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Principles of physically assisted extractions and applications 71
3 Principles of physically assisted extractions and applications in the food, beverage and nutraceutical industries E. Vorobiev, Compiègne University of Technology, France and F. Chemat, University of Avignon and Pays de Vaucluse, France Abstract: The physical methods used to enhance pressure and solute extraction methods in the food industry are reviewed. The quality of extracted products (e.g. purity, colour, texture, flavour and nutrients) may be degraded by conventional mechanical, thermal or chemical pretreatments. Several physical treatments (pulsed electric fields, power ultrasound, microwaves, ohmic heating, arc discharges) have been shown to be important and of great interest for the food industry in terms of enhancing solute and pressure extraction, and dehydration processes. Emerging extraction technologies that show promise for commercial food processing are discussed. Key words: plants, food, extraction, pulsed electric field, ultrasonic extraction, microwave extraction.
3.1 Introduction Pressure and solvent extractions are widely used in the food industry for the production of juices, wine, sugar and vegetable oil; these methods are also frequently applied in the extraction of various target compounds, such as colorants, antioxidants, essential oils, and aromas, from raw plant materials. Pressing, hot water and organic solvent extractions are long established processes that have excellent efficacy when applied in optimal conditions. These extractions can be done concurrently but more often they compliment each other and are technologically combined (e.g. the hot water extraction of sugar from sugar beets is combined with the subsequent pressing of pulps; the pressing of oilseeds is combined with the subsequent solvent extraction of oil from press-cake; the pressing of apples or grapes can be combined © Woodhead Publishing Limited, 2010
72 Separation, extraction and concentration processes with the subsequent solvent extraction of bioactive compounds from mash). The yield of extracted compounds can be very high in optimal conditions. Unfortunately, the quality of solutions and extracted products (e.g. purity, turbidity, colour, texture flavour and nutrients) may be degraded in the course of the raw material treatments that are necessary to increase the yield (grinding, heating, chemicals/enzyme addition). Moreover, a significant quantity of waste is often produced from the purification of impure extraction solutions, when undesirable loss of solvents and other additives occurs. These problems need to be resolved or minimised for the future development of environmentally sustainable food technologies. In recent decades there has been a growing interest in alternative food technologies that allow for nonthermal or mild thermal food preservation. Several emerging technologies are prominent and of great interest to the food industry, in particular pulsed electric fields (PEF), power ultrasound (PU), microwave (MW), pulsed light (PL), ohmic heating (OH), irradiation (IR), radio-frequency (RF) heating, high pressure processing (HP), high voltage electric discharges (HVED), amongst others (Barbosa-Cánovas et al., 1998; Zeuthen and Bogh-Sorensen, 2000; Povey and Mason, 1998). Such success in the development of novel preservation technologies has encouraged further research and has renewed the industrial interest in extractions assisted by nonthermal or mild thermal physical treatments (PU, PEF, HVED, OH, MW). Recently, these treatments were found to be effective for the enhancement of solute and pressure extractions, and dehydration processes (Knorr et al., 2001; Kulshrestha and Sastry, 2003; Li et al., 2004; Virot et al., 2007; Vorobiev and Lebovka, 2006, 2008). The potential benefits of physically assisted extractions are important. For instance, they can lead to the future substitution of hot water or organic solvent extractions with cold, or mild thermal, aqueous or pressure extractions. Moreover, the alternative physical treatments are found to be less invasive methods for the processing of plant foods, making it possible to avoid many undesirable changes in products, pigments, vitamins, and flavouring agents, which are typical of other extraction techniques, including thermal, chemical and enzymatic techniques. Some of these novel technologies remain very much in the research arena, whereas some others are on the brink of commercialisation. This chapter provides an overview of the emerging extraction technologies that are promising in terms of their potential applications in commercial food processing.
3.2 Pulsed electric field-assisted extractions in the food industry 3.2.1 Principles of pulsed electric field (PEF) treatment Pulsed electric field processing is a technique in which a food is placed between two electrodes in a batch or continuous treatment chamber and © Woodhead Publishing Limited, 2010
Principles of physically assisted extractions and applications 73 exposed to a pulsed voltage (typically 15–80 kV cm–1 with several pulses of 1–5 ms for the killing of micro-organisms; and 0.1–5 kV cm–1 with pulses of 50–1000 ms for the electroporation of plant cells and nonthermal extraction from solid foods). To generate such short pulses, various pulse-forming networks are used, with main components that include power supply at the selected voltage, one or several capacitor banks, inductors and/or resistors (Barbosa-Cánovas et al., 1998; Bluhm, 2006). Various different pulse shapes can be generated, including the simplest exponential decay pulses and squarewave pulses (mono-and bipolar). The duration and/or number of pulses are limited in order to prevent any significant temperature elevation, which is usually less than 3–5 °C. Mechanism of cell membrane electroporation When the intensity of an applied electric field increases, the potential difference across a cell membrane (transmembrane potential um) also increases. If um exceeds a stated threshold value (typically 0.2–1 V), a temporary loss of membrane semipermeability occurs. This phenomenon of membrane damage is named electroporation (electropermeabilisation) (Fig. 3.1). A number of models have been proposed in the biophysical literature to describe the electroporation mechanism (Konduser and Miklavcic, 2008; Zimmermann,
E < Ecr
Compression
Compressed membrane
E > Ecr
(a) Repulsion
Pore
Membrane
2r
(b)
Fig. 3.1 Membrane electroporation. With application of electric field E, the cell membrane is compressed owing to the electric charges attraction (a). (b) When the electric field exceeds some critical value (E > Ecr), the transmembrane potential increases to a threshold value and electroporation occurs (from Vorobiev et al., 2005).
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74 Separation, extraction and concentration processes 1986). The transmembrane potential um of a spherical cell depends on the angle q between the external field E direction and the normal vector on the membrane surface. The transmembrane potential may be given by the following equation (Zimmermann, 1986):
um = 0.75fdcEcosq [1 – exp (– t/tc)]
[3.1]
where dc is the cell diameter, tc is the time of membrane charging, and f is a parameter depending on the electrophysical and dimensional properties of the cell, membrane, and surrounding media. The values of f and tc vary for different cell sizes and ratios of extracellular/intracellular conductivities se/ si (Lebovka et al., 2000; Vorobiev and Lebovka, 2008). At complete cell damage, it follows that se/si = 1 and f ≈ 1. The critical value of transmembrane potential required for biological membrane electroporation at ambient temperature is estimated as um ≈ 1 V (Zimmerman, 1986) and can vary in plant and animal cells from 0.7 to 2.2 V (Knorr et al., 2001). The typical size of a cell in plant tissue is dc ≈ 50–100 mm. Therefore, the required threshold PEF intensity for the electroporation of food plants tissues (for f = 1 and tc <
Z = (s – si)/(sd – si)
[3.2]
where s is the electrical conductivity value measured at low frequency (1–5 kHz) and subscripts ‘i’ and ‘d’ refer to the conductivities of intact and totally damaged cells, respectively. This equation gives Z = 0 for intact tissue and Z = 1 for totally disintegrated material. Another method is based on electrical conductivity measurements at low (ª1 kHz) and high (3–50 MHz) frequencies (Angersbach et al., 2002). Figure 3.2 shows the characteristic time tE that is necessary for halfdamage (i.e. Z = 0.5) of red beet tissue (Shynkaryk et al., 2008). Inserts show examples of the conductivity disintegration index Z as a function of
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Principles of physically assisted extractions and applications 75 Conductivity disintegration index, Z
101
100
Time, tE (s)
10–1
10–2
1
E = 600 V cm–1
400 V cm–1 300 V cm–1 200 V cm–1
0.5
tE
10–5
10–4 10–3
10–2 10–1 tt (s)
100
101
10–3
10–4
10–5
200
300 400 500 600 Electric field strength, E (V cm–1)
700
Fig. 3.2 Characteristic time of half-damage (Z = 0.5) of red beet tissue (Shynkaryk et al., 2008), with insert showing examples of the conductivity disintegration index Z as a function of PEF treatment time.
treatment time tt = nNti, where ti = 100 ms is the duration of the square pulse, n is the number of pulses, and N is the number of trains. The PEF treatment was conducted at a room temperature, T = 20 °C. The value of tE decreases with the increase of the electric field strength E and its approaching is tE ª10–4 s at E ≥ 600 V cm–1. The compression-to-failure and the stress–relaxation measurements for apple, carrot and potato tissues treated using PEF over different time periods tPEF were presented by Lebovka et al. (2004b). After undergoing PEF treatment of rather high intensity and long duration (E = 1.1 kV cm–1, tPEF = 0.1 s), the tissues had lost part of their initial strength. However, the changes of both the elasticity modulus Gm and the fracture stress PF were significantly smaller than the changes observed for the freeze-thawed and thermally (T = 45 °C, 2 h) pre-treated tissues. It was concluded that PEF enables a high disintegration of membranes to be achieved, and a turgor component to be removed from the texture. However, compared with tissue treated by freeze-thawing or heating, the tissue structure seems to be less affected by PEF treatment. This conclusion was later confirmed by the textural studies of sugar beet tissue treated by PEF (Shynkaryk et al., 2008). Microscopic studies of onion tissue treated by PEF have shown intact cell architecture, whereas colorant could easily diffuse inside of cells with electroporated membranes (Fincan and Dejmek, 2002). Lebovka et al. (2000, 2001) have put forward a hypothesis on how PEF
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76 Separation, extraction and concentration processes treatment affects cellular tissues. They consider the effect to be correlated percolation which is governed by two processes. The first of these is the resealing of cells, and the other is the transfer of moisture inside the cellular structure, which is sensitive to repetition of PEF treatments. When the treatment is applied using a low enough intensity of electric fields, the electroporation is reversible as the resealing process is quick to repair the membranes immediately after the PEF treatment has been terminated. At moderate PEF treatment, some of the cells lose their permeability, but in others the pores may persist (Lebovka et al., 2001). High-intensity PEF treatment causes irreversible damage to the cell membrane. Long-term changes in conductivity after the application of PEF treatment can also be related to osmotic flow and to the redistribution of moisture inside the sample (Lebovka et al., 2001). PEF-treatment chambers Special chambers (batch or continuous) have been developed for PEF-assisted extraction (Bluhm and Sack, 2008; Jaeger et al., 2008; Lazarenko et al., 1977; Vorobiev and Lebovka, 2008). The chambers specified for the treatment of liquid foods by high PEF (Barbosa-Cánovas et al., 1998) can also be used for the pretreatment of yeasts before extraction (Shynkaryk et al., 2008). The chambers developed for the sterilisation of particulate foods by ohmic heating (Biss et al., 1989) can be adapted for the PEF pretreatment applied before the extraction of apple mash (Jaeger et al., 2008). In some cases, the PEF treatment can be combined with extraction in the same apparatus (Vorobiev and Lebovka, 2008). Figure 3.3(a) and (b) presents the laboratory treatment cell developed in Compiegne University of Technology, which combines PEF treatment and juice expression from solid foods. 3.2.2 Juice expression and solute extraction from plant roots, tubers and fruits In raw food plants, the valuable compounds are initially enclosed in cells, which have to be damaged to facilitate the recovery of intracellular matter. Conventional techniques of cell damage, such as fine mechanical fragmentation, and thermal, chemical and enzymatic treatments, lead to the more severe disintegration of tissue components, including cell walls and cell membranes. PEF treatment, which is less destructive than conventional methods, can be used for the more selective extraction of cell components. Sugar beets The conventional extraction technology consists of a power-consuming hot water diffusion of sugar from sliced cossettes at 70–75 °C. Unfortunately, the denaturation of tissue by heat causes alteration in the cell wall structure through hydrolytic degradation reactions. In addition to sucrose, other cell components, such as pectin pass in juice during extraction, affecting the juice
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Principles of physically assisted extractions and applications 77 PEF generator
Data processing
Compressed air
Steel frame Elastic rubber Pressure
Polypropylene ring Mobile electrode Sample Stationary wire gauze electrodes
Rubber shims
Filter cloth
Steel cover Juice treatment cell (a)
(b)
Fig. 3.3 Laboratory cell combining PEF treatment and juice expression from solid foods: (a) schematic diagram; (b) photographs.
purity. Consequently, a complex multi-staged juice purification process is used in beet sugar production (Van der Poel et al., 1998). The PEF application has great potential as an alternative method to the conventional thermal
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78 Separation, extraction and concentration processes technology. Recently, PEF-assisted pressing and aqueous extraction from sugar beets have been intensively studied (Bouzrara and Vorobiev, 2000; Bouzrara, 2001; El-Belghiti and Vorobiev, 2004, 2005a; Eshtiaghi and Knorr, 2002; Jemai and Vorobiev, 2003, 2006; Lebovka et al., 2007). Several studies have demonstrated the efficiency of PEF treatment for the cold pressing of sugar beet cossettes (Bouzrara and Vorobiev 2000; Bouzrara, 2001; Eshtiaghi and Knorr, 2002). Bouzrara and Vorobiev (2000) and Bouzrara (2001) thus reported that up to around 82% of the overall yield could be achieved by two-stage pressing with intermediately applied PEF (E = 400 V cm–1, tCEP = 0.1 s). The initial pressurisation of slices serves to assure a good electrical contact between them. In addition, the secondary juices (i.e. after PEF application) were systematically more concentrated in sugar and had lower colour concentration. PEF treatment was successfully applied with scale-up experiments exploring the viability of a novel process of cold juice extraction (Jemai and Vorobiev 2006). The processing scheme consisted of two initial pressing steps with an intermediate PEF treatment, followed by one or more washing steps and a final pulp pressing. The cold juices expressed from sugar beet gratings after the intermediate PEF treatment have higher purity values (95~98%) when compared with those before PEF application (90~93%). Additionally, the quantity of pectin was noticeably lower and the colour concentration of the juice was systematically 3 to 4 times lower than the colour of factory juices (Jemai and Vorobiev, 2006). These results, which show significant amelioration of the qualitative juice characteristics, give interesting new perspectives on cold PEF-enhanced expression from the sugar beets. Figure 3.4 shows a schematic diagram and a photograph of a pilot belt press that has recently been used for the PEF-assisted expression from sugar beets (Grimi et al., 2008). The results presented in Fig. 3.4 confirm the amelioration of juice yield (Fig. 3.4c) and purity (Fig. 3.4d) with the application of PEF pre-treatment. The size of particles treated by PEF should be optimised to obtain the juice with a maximal yield and higher purity level. Alternative studies have demonstrated the possibility of sugar extraction in cold or moderately heated water (Jemai and Vorobiev, 2003). Lebovka et al. (2007) compared the kinetics of thermal and cold diffusion and diffusion coefficients Deff for untreated and PEF-treated (E = 400 V cm–1, tCEP = 0.1 s) sugar beet slices (1.5 mm ¥ 10 mm ¥ 10 mm). The difference in Deff values between the untreated and PEF-pretreated tissue increases significantly for less heated tissue. For instance, the values of Deff were nearly the same for sugar diffusion at 60 °C from untreated tissue and at 30 °C from PEFpretreated tissue. The purest juice was obtained after cold diffusion. However, even after thermal diffusion at 70 °C, the juice purity was higher in slices pretreated by PEF than for untreated slices. Encouraging results obtained by several research groups revealed strong industrial interest, and semi-industrial scale equipment was built for PEFassisted extraction from sugar beets (Bluhm and Sack, 2008).
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Principles of physically assisted extractions and applications 79 Other roots and tubers In red beetroot, the water-soluble betalaines (red-violet betacyanins and yellow betaxanthins) are the main red–purple pigments. The high degree of extractability from red beetroot was observed after the PEF treatment at field strength 1 kV cm–1, when samples released about 90% of the total red
Upper belt F
C
J Drainage area
Low belt
J (a)
(b)
Fig. 3.4 Pilot belt press: (a) schematic diagram (F, feeding; C, cake; J, juice extract); (b) photograph; (c) juice yield; and (d) juice purity (from untreated and PEF-treated sugar beet slices of different sizes: S1 0.045 mm3, S2 47.5 mm3, S3 280 mm3 and S4 1050 mm3).
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80 Separation, extraction and concentration processes 100 Untreated PEF treated
Juice yield (%)
80
60
40
20 0
S 1
<
S 2
< (c)
S 3
<
S4
100
Purity (%)
Untreated PEF treated
95 90 85 80
S 1
<
S 2
< (d)
S 3
<
S4
Fig. 3.4 Continued
pigment following 1 h of aqueous extraction (Fincan et al., 2004). Lopez et al. (2009a) have shown that the application of PEF treatment at 7 kV cm–1 facilitates an increase in the maximum yield of betanine by a factor of 4.2, compared with PEF treated samples, achieving almost complete betanine release. The combination of PEF at 7 kV cm–1 and pressing at 14 kg cm–2 shortened the extraction time 18-fold. It can be noted that the effective electroporation of red beet tissue at ambient temperature can be attained even at the lower electric fields strengths of 400–600 V cm–1 (Shynkaryk et al., 2008; Fig. 3.2). For carrots, the PEF-assisted expression (Bouzrara, 2001; Knorr et al., 1994; Praporscic et al., 2007b), aqueous extraction (El-Belghiti and Vorobiev, 2005b) and their combination (Grimi et al., 2007) were studied. A significant enhancement of juice yield can be attained, even at a rather low voltage gradient of 360 V cm–1 (Bouzrara, 2001). Juices expressed with PEF treatment were more transparent and less turbid than untreated juices. Moreover, the Brix values were instantly increased after the PEF application (Praporscic et al., 2007b). Grimi et al. (2007) studied carrot juice extraction using a laboratory filter-press chamber with various combinations of pressing © Woodhead Publishing Limited, 2010
Principles of physically assisted extractions and applications 81 and washing operations. They have shown that it is possible to produce from the press-cake a ‘sugar-free’ concentrate, rich in vitamins and carotenoids, that can be used as an additive in diet foods. Potato tissue can also be effectively electroporated by PEF. Its textural and compressive properties after PEF treatment were studied by Lebovka et al. (2004b, 2007), and Grimi et al. (2009a). A prototype for potato starch extraction by PEF was developed by ProPuls, Germany, and commercial prototypes for this application based on a patented Marx Generator design have been developed by KEA-Tec, Germany (Jaeger et al., 2008). Apples Various research groups have studied the influence of PEF treatment on the expression of apple juice (Bazhal and Vorobiev, 2000; Bazhal et al., 2001; Lebovka et al., 2004a; McLellan et al., 1991; Praporscic et al., 2007b; Schilling et al., 2007). The reported improvement of juice yield varied, probably owing to the different processing conditions employed in the different studies (degree of particle fragmentation, PEF parameters, and compression pressure). For instance, the size of particles and fragmentation method (slicing, grinding or milling) can be essential for the improvement of juice yield. Recently, Grimi et al. (2008) demonstrated that the juice yield obtained from Golden Delicious apple slices (2 ¥ 3.5 ¥ 55 mm) after PEF treatment (E = 400 V cm–1, tCEP = 0.1 s) increased by 28%, whereas it increased by just 5% in finer slices (1 ¥ 1.9 ¥ 55 mm). The finely fragmented particles cause most cells to be disrupted mechanically, and the additional effect of the PEF on the total juice yield is rather limited. On the contrary, with coarse particles, the percentage of cell membranes damaged electrically increases, but the cell wall structure is less affected. This might be the reason for the more transparent and less cloudy apple juices obtained after PEF treatment of coarse particles (Bazhal and Vorobiev, 2000; Praporscic et al., 2007b). No apparent change in pH value or total acidity was detected. Additionally, the content of many nutritionally valuable compounds was retained, or even enhanced (Jaeger et al., 2008).The good industrial potential for PEF-assisted apple juice expression is confirmed in the laboratory and pilot scales using a belt-press equipment (Grimi et al., 2008; Jaeger et al., 2008). In addition, Jemai and Vorobiev (2002) have demonstrated that PEF exerts an enhancing effect on the diffusion of soluble substances from apple tissue. Apple discs (Golden Delicious) were pretreated thermally (75 °C, 2 min) and by PEF (E = 100–500 V cm–1, tPEF = 0.1 s), and afterwards the diffusion kinetics was studied at various temperatures. Experiments revealed that a detectable enhancement of diffusion kinetics starts at field intensities of 100–150 V cm–1. A further increase of both the field intensity and the pulse duration led to an enhancement of the diffusion kinetics. Jemai and Vorobiev (2002) showed that for thermally treated samples, the temperature variation of the diffusion coefficient D is of Arrhenius type with two diffusion
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82 Separation, extraction and concentration processes regimes: (i) without thermal pre-treatment (Ea~28 kJ mole–1) and (ii) after thermal denaturation (Ea~13 kJ mole–1). Only one regime with intermediate activation energy (Ea~20 kJ mole–1) was observed for electrically treated samples. Grapes Electroporation of wine grapes is an alternative nonthermal process leading to a prudent extraction of colours and valuable constituents. Recently, Praporscic et al. (2007a) investigated quantitative (juice yield) and qualitative (absorbance and turbidity) characteristics of juices obtained during the expression of white grapes (Muscadelle, Sauvignon and Semillon). The experiments were carried out at an expression pressure of 5 bar using a laboratory compression chamber equipped with a PEF treatment system. PEF treatment with a field strength of E = 750 V cm–1 and the total treatment duration tPEF = 0.3 s was applied. The PEF treatment resulted in an increase of the final juice yield Yf of up to 73–78% compared with Yf ª 49–54% for the untreated grapes. A rather noticeable decrease of absorbance and turbidity was observed as a result of the PEF treatment for all the white grape varieties studied. Furthermore, Grimi et al. (2009b) have shown that PEF treatment enhances compression kinetics and extraction of polyphenols from Chardonnay grapes. Lopez et al. (2009b) have studied the application of a PEF treatment (5 kV cm–1, 50 pulses) to the de-stemmed, crushed and slightly compressed grape pomace (skins, pulp and seeds) of Cabernet Sauvignon grapes. The application of a PEF treatment to the pomace before the vinification process led to freshly fermented wines that were richer in colour intensity, contained more anthocyanins and tannins, and showed better visual characteristics. The PEF treatment has facilitated a reduction in maceration time during the vinification of Cabernet Sauvignon grapes from 268 to 72 h. A specially designed mobile facility was built to produce electric fields of up to 60 kV cm–1 in the reactor at a repetition rate of up to 15 Hz. It is able to handle a throughput of up to 1 t h–1 with an energy consumption of around 15 kWh t–1. Both red and white wine grapes have been treated (Bluhm and Sack, 2008). Yeast cells The disruption of yeast cells (Saccharomyces cerevisiae) is a very important step in the industrial extraction of bioproducts, such as valuable proteins, cytoplasmic enzymes, and polysaccharides), which are present inside the cells. A thermal treatment at T > 50 °C results in damage to the yeast membranes, but also causes the denaturation and degradation of many valuable intracellular components. Electroporation exerts a minimal undesirable impact on liquid components inside and outside the cell, and can be done without significant temperature increase and cell debris formation. The commonly reported values of field strength E needed for the disintegration of membranes in S. cerevisiae cells, are rather high (E > 7.5 kV cm–1) (Schrive et al., 2006;
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Principles of physically assisted extractions and applications 83 Zhang et al., 1994), though smaller fields can also affect the structure of cells. For instance, the early stages of damage in the yeast cells were observed at E < 7.5 kV cm–1, in a PEF treatment of long duration (El Zakhem et al., 2006a,b). The electroextraction of proteins from suspensions of S. cerevisiae was also observed at 3.2 kV cm–1 (Ganeva et al., 2003), but high efficiency (yield of 85%) required long extraction following PEF treatment (>4 h at 30 °C). Recently, Shynkaryk et al. (2009) demonstrated that a rather long application of PEF treatment (1000 trains, 200 pulses of 100 ms in each train) was needed to result in a high level of membrane disintegration (Z > 0.8), even at E = 10 kV cm–1. Moreover, the quantity of released high molecular weight intracellular components extracted with PEF was lower than that extracted with other methods (HVED and high-pressure homogenisation) (Shynkaryk et al., 2009).
3.3 Ohmic heating-assisted extractions in the food industry 3.3.1 Principles of ohmic heating (OH) treatment Ohmic heating (OH) processing is a technique in which a food is placed between two electrodes in a batch or continuous treatment chamber and exposed to a dc, ac or pulsed voltage (typically 20–80 V cm–1) to heat the food and kill any micro-organisms. OH generators, which use ac and highfrequency electric fields permit reduced electrolysis and product contamination, compared with dc designs (Lima et al., 1999; Sastry, 2005). The basic principle of OH is the dissipation of electrical energy in the form of heat, resulting in the generation of internal energy (Sastry, 2005):
q = sE2
[3.3]
where q is the heat power per unit volume, s is the electrical conductivity of food and E is the electric field intensity. OH processing is currently commercialised for the pasteurisation and sterilisation of particulate foods (Biss et al., 1989; Sastry, 2005). However, the effects of OH are not purely thermal. Several studies have noted that electropermeabilisation is also induced by OH (Kemp and Fryer, 2007; Kulshrestha and Sastry, 2003; Lima et al., 2001; Praporscic et al., 2006; Schreier et al., 1993). Lebovka et al. (2005) have studied the membrane damage of ohmically heated potato and apple tissues. The OH experiments conducted under ac treatment were started from room temperature T = 22 °C at different electric field strengths E, and the treatment was stopped when the temperature reached T = Tf = 50 °C. The total treatment time te was higher with lower E (Fig. 3.5), but the samples were not kept in the warm juice (at 40–50 °C) for more than 15 min. The thermal damage and softening of potato and apple tissues at such mild thermal treatment conditions can be neglected (Lebovka et al., 2004a; 2004b). But the conductivity disintegration index Z © Woodhead Publishing Limited, 2010
84 Separation, extraction and concentration processes increased considerably at the electric fields E > 20 V cm–1, both for apples and potatoes (Fig. 3.5), and these data provide evidence for the significance of electrically induced tissue damage. More recently, Lebovka et al. (2007) estimated the variation of Z values of sugar beet tissue during OH. They obtained electrical conductivity graphs for intact, maximally damaged and ohmically treated tissues at various temperatures and calculated Z(T) from equation [3.2] for various electric fields E. The results provide evidence for the significance of the electroporation effect in OH. This effect becomes more significant as E is increased from 40 to 100 V cm–1. Imai et al. (1995) have found that the OH behaviour and impedance values of Japanese white radish depended on the frequency of the electric field, which varied from 50 Hz to 10 kHz. At a constant voltage of 40 V cm–1, the time required to heat the sample to the desired temperature was longer for higher frequencies of electric field. It was concluded that the initial rapid heating of white radish at low frequency (50 Hz) is caused by the electroporation effect. 3.3.2 Juice expression and solutes extraction enhanced by OH The first studies of extractions enhanced by OH were completed in the former Soviet Union, and summarised by Rogov and Gorbatov (1974) and by Lazarenko et al. (1977). In Soviet publications, electroporation was known as ‘electroplasmolysis’, and the combined thermal and electric effects caused by OH were called ‘thermo-electroplasmolysis’ or ‘low-gradient ac, Tf = 50 °C
Potato Apple 103
0.6
te (s)
Conductivity disintegration index, Z
0.8
0.4 102
0.2
0
0
20
40 60 80 Electric field strength, E (V cm–1)
100
101
Fig. 3.5 The conductivity disintegration index Z and time of treatment te versus electric field strength E for potato and apple tissues ohmically heated to 50 °C by ac treatment.
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Principles of physically assisted extractions and applications 85 plasmolysis’ to underline the relatively low electric fields (typically 10–100 V cm–1) used for OH. Research interest in extractions assisted by OH processing was renewed in Europe and the USA during the 1990s and 2000s (Sastry, 2005; Schreier et al., 1993; Wang and Sastry, 2002) probably due to the successful commercialisation of OH systems for food sterilisation (Biss et al., 1989). This interest in OH extractions focused on the synergy of electrical and thermal effects on cell tissue, and the lower temperatures needed for effective membrane damage compared with conventional heating. On the other hand, the electrical field applied in OH-assisted extraction is lower than that for nonthermal PEF treatment. Therefore, electrical equipment can be simplified to deliver a lower output voltage than typical high voltage PEF generators. Various studies have recently demonstrated the effectiveness of OH pretreatment for the intensification of pressing and solute extraction (Kemp and Fryer, 2007; Kulshrestha and Sastry, 2003; Lima et al., 1999; Praporscic et al., 2005, 2006; Sensoy and Sastry, 2004; Wang and Sastry, 2002). For instance, Lima et al. (1999) ohmically heated apple cuts up to 40 °C with a voltage of E = 40 V cm–1 and wave frequencies of 4 and 60 Hz. Subsequent pressing at a constant rate (3 cm min–1) resulted in a higher juice yield than that obtained for untreated cuts, and the 4 Hz pretreatment was more effective than the 60 Hz pretreatment. More recently, Wang and Sastry (2002) have demonstrated that OH pretreatments of 40 or 50 °C by ac with 40 V cm–1 and 60 Hz enhanced the juice yield from apple samples. Compared with untreated samples, or those preheated by microwaves, the juices extracted from ohmically preheated apple tissue were visually similar in quality, and the yields were higher. The lower frequency of OH compared with MW heating may be the important factor that caused the greater increase in juice yield (Wang and Sastry, 2002). There may be other mechanisms involved in intra- and extracellular moisture flow during OH. In addition to its thermal effects and electrical breakdown, electro-osmosis may also be the reason for moisture flowing inside of cell tissue. Bazhal and Vorobiev (2000) have demonstrated that electro-osmotic flow was induced in apple tissue by a dc electric field of low voltage, which resulted in a significant increase in juice yield, even at ambient temperature. Solute extraction can also be enhanced by OH pretreatment. Kemp and Fryer (2007), Kulshrestha and Sastry (2003), and Schreier et al. (1993), showed that there is an increased diffusion of betanine dye from red beets treated by OH. In their interesting study, Kulshrestha and Sastry (2003) treated red beet cuts at 45 °C over a range of voltages from 0 to 23.9 V cm–1, and at frequencies of 0 (dc), 10, 50, 250 and 5000 Hz. As the electric field frequency decreased, the betanine content in the solution increased. This was explained by the increased time for which the cell membrane was charged at lower field frequencies. An exception to this tendency was detected for those tissues treated by dc, which was less effective than OH treatment at 10 Hz. All the mechanisms involved in tissue modification by
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86 Separation, extraction and concentration processes OH treatment have not yet been elucidated. Electrolysis phenomena might also be important in terms of process optimisation. Other fresh materials, for example, mint leaf (Sensoy and Sastry, 2004), potato (Praporscic et al., 2006), and sugar beets (Praporscic et al., 2005) can also be processed by OH-assisted extraction.
3.4 Extraction assisted by high-voltage electrical discharges and applications in the food industry 3.4.1 Principles of underwater high-voltage electrical discharges (HVED) Underwater (electrohydraulic) HVED is a technique in which solid particles are placed in a dielectric liquid (typically tap water) inside a chamber, containing a high-voltage (HV) needle electrode and a plated grounded electrode, and exposed to pulsed shockwaves (typically 40–60 kV cm–1, 2–5 ms) to produce a liquid breakdown and particle fragmentation. This technique has various applications (defragmentation of solid materials, such as concrete, treatment of municipal solid waste, and drilling) (Bluhm, 2006). Some publications propose the use of the HVED application for the deactivation of micro-organisms and for the enhancement of extractions (Grémy-Gros et al., 2008). Figure 3.6 shows a schematic diagram of a HVED generator, as well as typical voltage and current curves measured during a HVED (Boussetta et al., 2009a). Effects of HVED The mechanisms of HVED are very complex and not yet well understood. The phenomenon is based on the electrical breakdown of water (Bluhm, 2006). Air bubbles that are already present in the water, or formed as a result of local heating, participate in and accelerate this phenomenon. If the electrical field is intense enough, the avalanche of electrons becomes a starting point for streamer propagation from the HV needle electrode to the grounded one (Fig. 3.7). A crack can be attracted from the discharge channel into the solid inclusion if their dielectric properties are different. The electrical breakdown is accompanied by a number of secondary phenomena (high-amplitude pressure shock waves, bubbles cavitation, and creation of liquid turbulence). For instance, the pressure exerted by the expanding channel almost always exceeds the tensile strength of the solid inclusions, and leads to the formation of cracks in the solids. Several hundred to thousands of discharges are sufficient to fragment the stones into small pieces of less than 1 mm (Bluhm, 2006). For biological particles, the cell structure can be damaged by HVED, which accelerates solute extraction (Grémy-Gros et al., 2008). If electrical breakdown does not occur, water carries electricity and behaves
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Principles of physically assisted extractions and applications 87 HVED generator
Vent hole
Insulation
Air Water + sample
Needle electrode Plate electrode
Metal
HVED treatment cell (a) 15
40
10
20 5
0
0
Current (kA)
Voltage (kV)
tf ª 0.5 ms
–5 0
2
4 6 Time (ms) (b)
8
10
Fig. 3.6 HVED treatment (a) experimental setup and (b) typical voltage and current graphs.
like a good electrical conductor; then electrical signals have exponential form. When HVED are produced, electrical signals look like damped oscillations, as presented in Fig. 3.6 (Boussetta et al., 2009a; Grémy-Gros et al., 2008). 3.4.2 Solute extraction enhanced by HVED HVED have been used to accelerate the extraction of solutes from soybeans, potato, tea leaves, peat, and fennel (Barskaya et al., 2000; El-Belghiti, 2005; El-Belghiti et al., 2007) as well as from other food materials.
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88 Separation, extraction and concentration processes
(a)
(b)
(c)
HV electrode
Shock wave
Charging unit
Ground Solid (d)
Fig. 3.7 Type of electric discharge in aqueous solution: (a) streamer, E = 4.4 kV cm–1; (b) streamer and spark, E = 13.3 kV cm–1; (c) spark, E = 33.3 kV cm–1 (Sugiarto et al., 2001). (d) Schematic diagram of electro-hydraulic disintegration process.
Oil The application of HVED can enhance the aqueous extraction of oil from oilseeds (Grémy-Gros et al., 2008; Gros et al., 2003). Linseeds were crushed and pressed with a hydraulic press, and then the press-cake obtained was reduced to powder and dissolved in demineralised water at ambient temperature. The mixture was treated with HVED (1–1640 discharges) and centrifuged (Gros et al., 2003). Approximately 26% of oil remained in the residue after 1640 pulses. The process was then optimised by varying the number of pulses, pH, water/press-cake ratio and temperature (Grémy-Gros et al., 2008). The HVED treatment also made it possible to enhance mucilage extraction from whole linseed (Gros et al., 2003). Seeds were dissolved in demineralised water and treated for 10 min (i.e. 300 pulses of 0.5 Hz frequency). A centrifuge separation was then performed to obtain solution and residue. The residue was then treated a second and a third time with fresh water under the same conditions. During the HVED treatment, linseeds were defragmented under the action of shock waves. Three successive 10-min treatments were sufficient to extract the mucilage almost entirely. The extraction of proteins began during the third treatment. However, the oil–protein–water emulsions obtained by HVED treatment then had to be separated. The following research study
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Principles of physically assisted extractions and applications 89 resulted in the development of membrane separation of emulsions created by HVED (Li et al., 2009). Grape pomace Grape pomace (composed of stems, seeds and skins) has a high polyphenol content. These polyphenols have attracted great interest as they exhibit antibacterial, antiviral, and antioxidant properties, and can help prevent cardiovascular diseases (Dugand, 1980). Boussetta et al. (2009b) studied the application of HVED in the extraction of polyphenols from the pomace obtained as a residue from pressed white grapes (Vitis vinifera L., cultivar ‘Chardonnay’, vintage 2007). The treatment chamber (Fig. 3.6) was initially filled with pomace mixed with distilled water (the liquid-to-pomace ratio was 3:1, the water temperature was 20, 40 or 60 °C), and 80 successive discharges were applied (40 kV, pulse repetition rate 0.5 Hz). Afterwards, the total solute and polyphenol diffusion from the untreated and treated pomace mixtures were studied at the desired temperatures. Figure 3.8 shows that both temperature and HVED treatment improved the extraction of polyphenols. The difference between the yield of solutes for experiments with HVED treatment and without HVED treatment decreased with temperature elevation (Boussetta et al., 2009b). On the contrary, the opposite tendency was observed for the yield of polyphenols. The application of HVED enhanced the polyphenol yield, not just for fresh grape pomace, but also for sulfured
1
20
T (°C) 40
60 With HVED treatment
0.8
Ypolyphenols (%)
0.6
0.4
Without HVED treatment 0.2
0.0034
0.0032 1/T (K)
0.003
Fig. 3.8 Effect of treatment temperature on final yield of polyphenols after 1 h of extraction of fresh grape pomace with and without HVED treatment.
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90 Separation, extraction and concentration processes and frozen pomaces. The final yields of solutes, reached after HVED and subsequent diffusion for 40 min were more than twice those reached after 240 min without HVED. Yeast cells Shynkaryk et al. (2008) studied the efficiency of HVED-assisted extraction from wine yeast cells (S. cerevisiae) in aqueous suspension (1% w/w). HVED induced damage in yeast membranes, leakage in cytoplasmic ions, and the release of intracellular bioproducts; it also resulted in an increase in the disintegration index Z. For HVED treatment at 40 kV, the high level of membrane disintegration (Z > 0.8) required more than 100 discharges, which corresponded to the effective time of HVED treatment between 100 and 200 ms. HVED treatment decreased the average cell size somewhat, owing to cell disruption. However, the quantity of high molecular weight intracellular components released, controlled by the absorbance analysis, was lower than that obtained by high-pressure homogenisation (Shynkaryk et al., 2008).
3.5 Ultrasound-assisted extraction (UAE) in the food industry 3.5.1 Principles of UAE Power ultrasound, which has frequencies between 20 kHz and 1 MHz, is now well known to have significant effects on the rate of various physical and chemical processes. Cleaning and solubilisation are its more developed applications, and a large variety of ultrasound baths exist for chemical laboratory use. The effect of ultrasonic waves on solid samples is widely used for the extraction of aromas from plant materials, or metal impurities from soils. Degassing and stripping are widely used for flavour analysis, and in environmental and polymer research. Other interesting ultrasound applications include homogenisation, emulsification, sieving, filtration, and crystallisation. The most interesting effect of ultrasound-based operational units is the reduction of processing time and the increase in product quality. All these effects are attributed to acoustic cavitation: when a liquid is irradiated by ultrasound, micro-bubbles form, grow and oscillate extremely fast, and eventually collapse powerfully (if the acoustic pressure is high enough). When the size of these bubbles reaches a critical point, they collapse during a compression cycle and release large amounts of energy. The temperature and pressure at the moment of collapse have been estimated to be up to 5000 K and 5000 atmospheres. This creates hotspots that are able to dramatically accelerate chemical reactivity in the medium. When these bubbles collapse onto the surface of a solid material, the high pressure and
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Principles of physically assisted extractions and applications 91 temperature released generate microjets directed towards the solid surface. These microjets are responsible for the degreasing effect of ultrasounds on metallic surfaces, which is widely used for cleaning materials. Another application of microjets in the food industry is the extraction of vegetal compounds (Mason, 1990; Suslick, 1988). As shown in Fig. 3.9, a cavitation bubble can be generated close to the surface of the plant material. Then, during a compression cycle, this bubble collapses and creates a microjet (400 km h–1) directed toward the plant matrix. The high pressure and temperature involved in this process will destroy the cell walls of the plant matrix, and its content can be released into the medium. Power ultrasound involves the mechanical and chemical effects of cavitation. The mechanism can be explained by two competing theories. The hot spot theory assumes that the high pressures and temperatures generated in the bubbles during the last nearly adiabatic compression, just before collapse, are responsible for the breakage of molecular bonds and the formation of radicals. On the other hand, the electrical theory describes micro-discharges, produced by the high electrical fields generated by the deformation and fragmentation of the bubbles. The most commonly used frequencies in sono-extraction are between 20 and 40 kHz. With higher frequencies, cavitation would be more difficult to induce because the cavitation bubbles need a slight delay to be initiated during the rarefaction cycle. The higher the frequency, the shorter the rarefaction cycle, so the less time available for the bubble to be created (Povey and Mason, 1998).
Cavitation bubble with negative pressure
Maximum and critical bubble size 2
1
A new cavitation bubble appears and cycle is repeated
3
4
Implosion of bubble by compression
Fig. 3.9 Ultrasonic cavitation phenomenon and its action on extraction of natural products.
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92 Separation, extraction and concentration processes 3.5.2 UAE reactors The two most common types of ultrasound equipment that are used for extraction are the ultrasonic cleaning bath and the more powerful probe system. For small extraction volumes, an ultrasound horn with its tip submerged in the fluid can be sufficient. Large volumes of fluids have to be sonicated in an ultrasound bath or in continuous or recycled-flow sono-reactors. There are also ultrasonic reactors conceived especially for solvent extraction on a laboratory scale (3 l), in a pilot plant (30 l) and on an industrial scale (150–1000 l). Pump systems are coupled to the ultrasonic reactor to stir the mixture and to empty the system at the end of the experiment. The intensity of the ultrasound is about 1 W cm–2, with a frequency of 25 kHz. In order to maintain constant temperature, the reactor is made of a double mantle into which cooling water can circulate. The main advantage of this type of apparatus is that the natural products and extraction solvent are mixed into a container, and the ultrasound is directly applied to the mixture (Fig. 3.10). Recently, the new methodology of continuous-flow systems has been used in analytical chemistry. Most UAE applications have been developed in discrete systems, using a bath or an ultrasonic probe, particularly in the extraction of food samples. Less frequent has been the design of on-line UAE systems in the same field. However, it is noted that the latter approach is considerably faster. It consists of an open system, in which fresh solvent flows continuously through the sample. This induces the displacement of mass transfer equilibria toward the solubilisation of analytes into the liquid phase (Priego-Capote and Luque de Castro, 2007). 3.5.3 Factors affecting UAE All the research carried out over the last 50 years strongly supports the importance of each identified factor in the ultrasound process, where frequency, intensity, treatment time, temperature, pressure, and treatment
(a)
(b)
Fig. 3.10 (a) Laboratory (3 l) and (b) industrial (150 l) ultrasound extraction reactors (with permission, www.etsreus.com).
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Principles of physically assisted extractions and applications 93 media define the extraction and the separation kinetics of ingredients from natural products (De Gennaro et al., 1999). Generally, the highest efficiency of UAE, in terms of the yield and composition of the extracts, can be achieved by increasing the ultrasound power, reducing the moisture of food matrices to enhance solvent–solid contact, and optimising the temperature to allow a shorter extraction time. Solvent choice is dictated by the solubility of the ingredients of interest, the interactions between the solvent and matrix, and the intensity of ultrasound cavitation phenomena in the solvent. The impact of ultrasound is attributed to intracellular cavitation, that is, micro-mechanical shocks that disrupt cellular structural and functional components up to the point of cell lysis (Butz and Tauscher, 2002). Figure 3.11 shows a proposed mechanism for ultrasound-induced cell damage. Cavitation is the formation, growth and, sometimes, the implosion of microbubbles created in a liquid when ultrasound waves propagate through it. The collapse of the bubbles leads to energy accumulation in hot spots, where temperatures of above 5000 °C and pressures of approximately 500 MPa have been measured. This phenomenon can cause enzyme inactivation through three mechanisms, which can act alone or combined. The first one is purely thermal, owing to the enormous temperatures achieved during cavitation. The second is caused by free radicals, generated by water sonolysis; and the third is the result of the mechanical forces (shear forces) created by microstreaming and shock waves (Raso et al., 1999). 3.5.4 UAE: main applications Essential oils and aromas UAE has also been developed for the extraction of essential oils from aromatic plants, such as peppermint leaves (Shotipruk et al., 2001), artemisia (Asfaw et al., 2005) and lavender (Da Porto et al., 2009), or from other vegetal matrices, such as garlic (Kimbaris et al., 2006) and citrus flowers (Alissandrakis et al., 2003). Increased yields of essential oil were found for peppermint leaves (up to 12%) and for artemisia when using UAE; and yields of the main compounds of lavandula essential oil were found to increase by Ultrasound (US)
US
H 2O
Cell
US
Cell
Pore initiation
US
H2O influx and swelling
Cell
Membrane rupture and cell lysis
Fig. 3.11 Mechanism of ultrasound-induced cell damage.
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94 Separation, extraction and concentration processes 2- to 3-fold when comparing UAE with conventional distillation. Moreover, UAE not only improved yields, but also showed less thermal degradation in the final extracts. Several studies have been conducted into the extraction of the main aroma compounds from spices. For example, vanillin was extracted from vanilla pods (Jadhav et al., 2009) carvone from caraway seeds (Chemat et al., 2004) and safranal from Greek saffron (Kanakis et al., 2004). Yields of vanillin obtained after 1 h using UAE were similar to those obtained after 8 h using conventional extraction methods. Wines and other alcoholic beverages contain many volatile compounds, which play an important role in organoleptic characteristics. The ultrasound-assisted maceration of fruits and herbs in alcoholic beverages to extract volatile compounds is an important process in the food industry (Fig. 3.12). Antioxidants A wide variety of fruits and vegetables have been studied by UAE because antioxidants are present in different amounts in different varieties of plants, and these antioxidants come from different families. One of the most common antioxidants is the lycopene extracted from tomatoes (Lianfu and Zelong, 2008). In this particular work, authors not only worked on UAE, but also on coupling ultrasounds with microwaves, which produces high lycopene extraction in only 6 min. Herrera and Luque de Castro (2005) showed that the amount of phenolic compounds that could be extracted from strawberries in 2 min, was similar to the amount extracted over 20 h using the conventional method, and 3 h using supercritical fluid extraction. Another example is the extraction of anthocyanins from raspberries, developed by Chen et al. (2007). In this work, the same yields of anthocyanins were achieved in 3 min when using UAE, compared with 53 min when using the conventional extraction system. The difference between ultrasound and conventional methods will be more or less significant, depending on the part of the plant studied. This has been highlighted by studies of grapes and grape seeds (Palma and Barroso, Sono-extraction
0.100
Conventional extraction
Yield (%)
0.075 0.050
0.025 0.000
0
30
60
90
120 150 180 210 240 270 Time (min)
Fig. 3.12 Rapid sono-extraction of orange peels in alcoholic beverages.
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Principles of physically assisted extractions and applications 95 2002). In this work, authors found that the use of ultrasound instead of maceration was more interesting when working on seeds than on the whole fruit (Table 3.1). Oil and fat Many papers have reported on the UAE of oil and fat from various food samples. According to Luque-Garcia and Luque de Castro (2004), extraction from oleaginous seeds is difficult. Indeed, only 75–85% of the oil is solubilised in the solvent. The rest of the oil content is strongly bound to the matrix and cannot be extracted without additional treatments. For instance, Li et al. (2004) described the UAE of soybean oil in an optimised yield and reduced operating time, compared with conventional maceration. Luque-Garcia and Luque de Castro (2004) have reported a device consisting of an ultrasoundassisted Soxhlet apparatus, which is an attractive alternative to traditional Soxhlet extraction. Using this device, it takes only 90 minutes to obtain the same yield (99% of fat recovery) as can be obtained over 12 h with the traditional Soxhlet extraction method (Table 3.1). 3.5.5 Hazard analysis critical control point (HACCP) for UAE processing operation The hazard analysis critical control point (HACCP) system is a process that identifies and assesses the hazards and risks associated with the manufacture, Table 3.1 Ultrasound-assisted extraction of food ingredients Matrix (reference)
Extracts
Essential oils and aromas Citrus flowers (Alissandrakis et al., 2003) Caraway seeds (Chemat et al., 2004) Garlic (Kimbaris et al., 2006) Lavender flowers (Da Porto et al., 2009) Greek saffron (Kanakis et al., 2004) Antioxidant extracts Strawberries (Herrera and Luque de Castro, 2005) Raspberries (Chen et al., 2007) Tomatoes (Liangfu and Zelong, 2008) Grape (Palma and Barroso, 2002) Fat and oil extraction Olive seeds (Luque-Garcia and Luque de Castro, 2004) Soybean germ and seaweed (Li et al., 2004) Cocoa powder and nibs (Luque- Garcia and Luque de Castro, 2004) Bakery products (Luque-Garcia and Luque de Castro, 2004)
Linalool Carvone and limonene Essential oil (organosulfur compounds) Essential oil (1,8-cineol, camphor and linalyl acetate) Safranal Polyphenolic compounds Anthocyanins Lycopene Phenolic Oil Oil Fat content Fat content
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96 Separation, extraction and concentration processes storage and distribution of foods and implement the appropriate controls, aiming at the elimination or reduction of these hazards at specific points of the production line. In UAE processing, parameters such as the temperature and physical properties of the product, the timescale of treatment in the ultrasound chamber, the frequency and power of the ultrasound, and the nature of the probe are points which need to be monitored. For each part of the food process, including ultrasound processing, the evaluation and classification of hazards are completed. In this way, a grade is given to each hazard that has been identified according to its gravity, its risk of occurrence and the ability to detect it. For each hazard, preventive measures are set up, with procedures that indicate who will be responsible for them, as well as how, when and where. Preventive measures are aimed primarily at avoiding the occurrence of the hazard. For example, a good maintenance plan decreases the risk of metal contamination from the ultrasound horn (Table 3.2).
3.6 Microwave-assisted extraction (MAE) in the food industry 3.6.1 Principles of MAE Microwaves are electromagnetic waves with a frequency range of between 100 MHz and 3 GHz that comprise electric and magnetic field components and thus constitute propagating electromagnetic energy. This energy acts as a non-ionising radiation, which causes molecular motion of ions and rotation of the dipoles, but does not affect molecular structure. Table 3.2 Possible critical control point limits and associated corrective actions in an ultrasound-assisted extraction processing Critical control point
Danger
Target
Deviation
Processing Temperature Adequate Does not conform operation to the requirements Ultrasound Adequate Does not conform frequency to the requirements Ultrasound Adequate Does not conform power to the requirements Flow rate Below Does not conform specifications to the requirements Probe Absence Detected spoilage
Corrective actions
Adjust temperature reject or reprocess the product Adjust frequency, laboratory controls, or reprocess the product Adjust the power, laboratory controls, or reprocess the product Adjust flow, laboratory controls, or reprocess the product Reject
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Principles of physically assisted extractions and applications 97 When dielectric materials containing either permanent or induced dipoles are placed in a MW field, the rotation of the dipoles in the alternating field produces heat. More precisely, the applied MW field causes the molecules to spend slightly more time, on average, orienting themselves in the direction of the electric field rather than in other directions. When the electric field is removed, thermal agitation returns the molecules to a disordered state in the relaxation time and thermal energy is released. Thus, MW heating results from the dissipation of the electromagnetic waves in the irradiated medium (Fig. 3.13). The amount of dissipated power in the medium depends on the complex permittivity of the material and the local time-averaged electric field strength (Metaxas and Meredith, 1993). In conventional heating, heat is transferred from the heating medium to the interior of the sample, whereas in MW heating; heat is dissipated volumetrically inside the irradiated medium, and thus heat transfers occur from the sample to the colder environment. This causes a significant difference Moisture (%)
Classical heating
Microwave heating
Microwave heating Heat transfer
1 2
1. Solid friction desorption 2. Internal diffusion 3. External diffusion
Mass transfer
3
Conventional heating Heat transfer Mass transfer
Fig. 3.13 Microwave versus conventional extraction.
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98 Separation, extraction and concentration processes between conventional and MW heating. In conventional heating, heat transfer depends on thermal conductivity, on the temperature difference across the sample, and, for fluids, on convection currents. As a result, the temperature increase is often rather slow. In contrast, the volumetric heating effect in MW heating enables much faster temperature increases to be obtained, depending on the MW power and the dielectric loss factor of the material being irradiated. The influence of MW energy on chemical or biochemical reactions is strictly thermal. The MW energy quantum is given by the standard equation W = hn. Within the frequency domain of microwaves and hyper-frequencies (300 MHz–300 GHz), the corresponding energies are respectively 1.24 ¥ 10–6 eV to 1.24 ¥ 10–3 eV. These energies are much lower than the usual ionisation energies of biological compounds (13.6 eV), of covalent bond energies like OH (5 eV), hydrogen bonds (2 eV), and van der Waals intermolecular interactions (lower than 2 eV) and even lower than the energy associated with Brownian motion at 37 °C (2.7 ¥ 10–3 eV). From a scientific point of view, the direct molecular activation of microwaves should be excluded. 3.6.2 MAE reactors It is now possible to use a wide range of vessels and instrumentation when working with microwaves, depending on the intended purpose and the analytes to be extracted. Two types of MW ovens can be found in laboratories and they differ in terms of their cavity type: single-mode (also called monomode or focused mode) and multi-mode. A single-mode cavity permits the generation of one type of frequency, allowing only one mode of resonance to be excited. Thus, the sample is placed in such a way as to receive the maximum amount of MW energy, ensuring a direct and efficient application of a high density of MW energy to the sample matrix, and inducing a rapid heating of the medium with a fast extraction of the analyte. However, monomode systems are quite limited in terms of the volume and quantity of the sample that is ultimately useable. On the other hand, multi-mode ovens are larger and allow several modes of resonance. Thus, a large number of incident waves can be applied to the sample being studied. This kind of oven allows for a large quantity of matrix, and can be used with several types and forms of vessels. There are also MW reactors conceived especially for solvent extraction in the laboratory (0.1 to 1 l) and on an industrial scale (100 l h–1) (Fig. 3.14). 3.6.3 MAE: main applications Microwaves are very useful for rapid extraction and obtaining high quality aromatic compounds from garlic, lavender flowers, orange peel and rosemary leaves. The extraction of essential oils from orange peels with solvent-free MW extraction is better in terms of energy saving, extraction time (30 min
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Principles of physically assisted extractions and applications 99
(a)
(b) –1
Fig. 3.14 (a) Laboratory (1 l) and (b) industrial (100 l h ) MAE reactors (with permission, www.milestonesrl.com and www.rchimex.com).
versus 3 h), product yield (0.42 versus 0.39%) and product quality (AbertVian et al., 2008). Hemwimon et al. (2007) made a comparison between the MAE of antioxidative anthraquinones from the roots of Morinda citrifolia, and extraction with other conventional techniques (maceration and Soxhlet). The efficiency of extraction using MAE (15 min) was much higher than that using maceration (3 days) and was also similar to that of Soxhlet (4 h). Similarly, the best recoveries of antioxidants in short time periods, and with a lower consumption of polar solvent were also observed from sweet grass by Grigonis et al. (2005). As well as obtaining a higher yield, Pan et al. (2008) extracted phenolic compounds with scavenging ability, compared with synthetic antioxidants from longan peel with MAE. Elkhori et al. (2007) described the MAE of cocoa powder with hexane/ isopropanol, which resulted in the rapid determination of fat contents, with recoveries similar to, or better than, those of the conventional method, and with the added advantages of low solvent consumption, short extraction time, low energy consumption and excellent reproducibility. Focused MW-assisted Soxhlet extraction was used by Pérez-Serradilla et al. (2007) for acorn oil determination, which resulted in trans fatty acid-free oil, probably because of the reduced exposure to drastic conditions in just 30 min, which is much less than the time required by the Soxhlet (8 h) and stirring (56 h) reference methods. Lianfu and Zelong (2008) described the extraction of lycopene from tomatoes with the combined innovative techniques of ultrasound and MW-assisted extraction (UMAE) in comparison with ultrasound-assisted
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100 Separation, extraction and concentration processes extraction (UAE). Similarly, a higher yield of anthocyanins in red raspberries was also observed by Sun et al. (2007) with optimal conditions of MAE. MAE proved to be more rapid and efficient as it extracts various types of anthocyanins, without any destruction of their chemical structure, in a very short time, owing to the intensive disruption of tissue structure under MW irradiation. Table 3.3 clearly shows some applications of MAE for various useful metabolites, carried out using several vegetal matrices. 3.6.4 Safety considerations The MAE process is simple and can be readily understood in terms of the operating steps to be performed. However, in inexperienced hands, the application of MW energy can pose serious hazards. All persons dealing with microwaves must exercise a high level of safety and attention to detail when planning and performing such experiments. They have to ensure that they seek proper information from knowledgeable sources and that they do not attempt to implement this technique unless proper guidance is provided. Only approved equipment and scientifically sound procedures should be used.
3.7 Combination of physical treatments for extraction in the food industry Various physical treatments can induce various effects on cell and tissue structure, which may be more or less beneficial for the extraction of intracellular components. Unfortunately, the studies comparing or combining two or more Table 3.3 Microwave-assisted extraction of food ingredients Matrix (reference)
Extracts
Antioxidant extracts Green tea leaves (Pan et al., 2003) Noni plant roots (Hemwimon et al., 2007) Sea buckthorn (Sharma et al., 2008) Longan peel (Pan et al., 2008)
Polyphenols Anthraquinones Phenolic constituents Flavonoids
Fat and oil extraction Olive seeds (Virot et al., 2007) Oleaginous seeds (Pérez-Serradilla et al., 2007) Soybean germ and seaweed (Cravotto et al., 2008) Cocoa powder and nibs (Elkhori et al., 2007)
Oil Oil Oil Fat content
Natural food colour extraction Tomato paste (Lianfu and Zelong, 2008) Curcuma (Mandal et al., 2008) Red raspberries (Sun et al., 2007) Paprika (Csiktusnádi Kiss et al., 2000)
Lycopene Curcumin Anthocyanins Carotenoids
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Principles of physically assisted extractions and applications 101 physically assisted extractions are rather rare. El-Belghiti (2005) compared the application of PEF (680 V cm–1, 1000 pulses of 100 ms) and HVED (40 kV, 100 discharges) for the cold (20 °C) aqueous extraction of solutes from two dried products: tea leaves and Datura innoxia (common name: moonflower) roots. PEF treatment did not have any influence on the process of extracting from dried and rehydrated products, probably because of the absence of the cell membrane element, which had been disrupted by previous drying. The non-efficacy of moderate electric treatment for the dried cell tissue was reported earlier in Sensoy and Sastry (2004). In contrast to PEF, the HVED smashed products to small fragments and the extraction kinetics were significantly enhanced for both tea leaves and Datura innoxia. El-Belghiti et al. (2007) compared PEF, HVED and PU as treatments to enhance cold (20 °C) aqueous extraction from fennel gratings. The objective was to obtain extracts used as natural food preservatives (antioxidants). The three assisted extractions led to the same final yield of solutes, and all preserved the antioxidant substances. However, their kinetics were different. The most rapid and the slowest kinetics were offered by HVED and PU, respectively. The final yield of 98% was reached in 20 min with HVED, in 40 min with PEF, and in more than 180 min with PU. The amounts of energy needed for these treatments were also different: the PEF treatment appeared to be the most economic energetically. Recently, Shynkaryk et al. (2008) compared PEF, HVED, and high-pressure homogenisation (HPH) for disruption of wine yeast cells (S. cerevisiae bayanus). The PEF and HVED were applied at electric field strengths of 10 kV cm–1 and 40 kV cm–1, respectively. The HPH was applied within a pressure range of 30–200 MPa for a different number of passes (1–20). It was shown that the releasing efficiency of ionic components, enzymes, proteins and other bioproducts dramatically depended on the applied method of disruption. Extraction efficiency was rather high for ionic components and small for high molecular weight components in PEF or HVED pretreated suspensions. The PEF treatment removes membrane barriers and accelerates the release of ionic contents, but has practically no influence on the cell walls. The efficiency of releasing high molecular weight content is questionable and may depend on the cell strain, the age of culture, the time of extraction, temperature and many other factors. For example, high levels of PEF extraction of proteins for baker’s yeast was previously reported (Ganeva et al., 2003), but the efficiency of the PEF extraction of proteins was found to be rather low for wine yeast in the study by Shynkaryk et al. (2008). On the contrary, the application of the combined HVED–HPH technique resulted in a material loss of integrity in the yeast cell walls. Even a small disintegration (Zi = 0.15±0.05), initially induced by HVED pretreatment, resulted in noticeable acceleration of HPH disruption kinetics. Although most of the research effort in UAE has concentrated on ultrasound itself, some studies have also examined the coupling between ultrasound and other techniques. For instance, UAE is currently being employed in
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102 Separation, extraction and concentration processes combination with MW energy (Lagha et al., 1999), supercritical fluid extraction (Hu, 2007) or simply with conventional methods such as Soxhlet extraction (Luque-García and Luque de Castro, 2004). When combined with supercritical fluid extraction, UAE enhances the mass transfer of the species of interest, from the solid phase to the solvent used for extraction. Soxhlet extraction can also be improved by ultrasound when applied at the cartridge zone before siphoning, thus permitting the removal of lipid fractions from very compact matrices. The efficiency of combining MW and ultrasound has been clearly shown in applications such as the extraction of food ingredients. Studies combining assisted extractions present undeniable interest as synergistic effects of different treatments may be discovered.
3.8 References Abert-Vian M, Fernandez X, Chemat F (2008), ‘Microwave hydrodiffusion and gravity, a new technique for extraction of essential oils’, Journal of Chromatography A, 1190, 14–18. Alissandrakis E, Daferera D, Tarantilis P A, Polissiou M and Harizanis P C (2003), ‘Ultrasound-assisted extraction of volatile compounds from citrus flowers and citrus honey’, Food Chemistry, 82, 575–579. Angersbach A, Heinz V and Knorr D (2002), ‘Evaluation of process-induced dimensional changes in the membrane structure of biological cells using impedance measurement’, Biotechnology Progress, 18(3), 597–603. Asfaw N, Licence P, Novitskii AA, Poliakoff M (2005), ‘Green chemistry in Ethiopia: the cleaner extraction of essential oils from Artemisia afra: a comparison of clean technology with conventional methodology’, Green Chemistry, 7, 352–356. Barbosa-Cánovas G V, Pothakamury U R, Palou E and Swanson B G (1998), Nonthermal preservation of foods, Marcel Dekker, New York. Barskaya A V, Kuretz B I and Lobanova G L (2000), ‘Extraction of water soluble matters from vegetative raw material by electrical pulsed discharges’, 1st International Congress on Radiation Physics, High Current Electronics, and Modification of Materials, Tomsk, Russia, 533–535. Bazhal M and Vorobiev E (2000), ‘Electrical treatment of apple cossettes for intensifying juice pressing’, Journal of the Science of Food and Agriculture, 80, 1668–1674. Bazhal M I, Lebovka NI and Vorobiev E I (2001), ‘Pulsed electric field treatment of apple tissue during compression for juice extraction’, Journal of Food Engineering 50, 129–139. Biss C H, Combes S A and Skudder P J (1989), ‘The development and application of ohmic heating for the continuous processing of particulate food stuffs’, In Processing engineering in the food industry (Field RW, Howell JA, eds), Elsevier, London, 17–27. Bluhm H (2006), Pulsed power systems, Springer, Berlin, Heidelberg, New York. Bluhm H and Sack M (2008), ‘Industrial-scale treatment of biological tissues with pulsed electric field’, In: E. Vorobiev and N. Lebovka (Eds), Electrotechnologies for extraction from food plants and biomaterials, Springer, 237–269. Boussetta N, Lebovka N, Vorobiev E, Adenier H, Bedel-Cloutour C, and Lanoisellé J-L (2009a), ‘Electrically assisted extraction of soluble matter from Chardonnay grape skins for polyphenols recovery’, Journal of Agricultural and Food Chemistry, 57, 1491–1497.
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Principles of physically assisted extractions and applications 103 Boussetta N, Lanoisellé J-L, Bedel-Cloutour C and Vorobiev E (2009b), ‘Extraction of polyphenols from grape pomace by high voltage electrical discharges: effect of sulphur dioxide, freezing process and temperature’ Journal of Food Engineering, 95, 192–198. Bouzrara H (2001), Amélioration du pressage de produits végétaux par Champ Electrique Pulsé. Cas de la betterave à sucre, PhD thesis, UTC, Compiègne, France. Bouzrara H and Vorobiev E (2000), ‘Beet juice extraction by pressing and pulsed electric fields’, International Sugar Journal, CII(1216), 194–200. Bouzrara H and Vorobiev E (2003), ‘Solid/liquid expression of cellular materials enhanced by pulsed electric field’, Chemical Engineering and Processing, 42, 249–257. Butz P and Tauscher B (2002), Emerging technologies: chemical aspects, Food Research International, 35, 279–284. Chemat S, Lagha A, AitAmar H, Bartels P, Chemat F (2004), ‘Comparison of conventional and ultrasound-assisted extraction of carvone and limonene from caraway seeds’, Flavour and fragrance journal, 19, 188–193. Chen F, Sun Y, Zhao G, Liao X, Hu X, Wu J, Wang Z (2007), ‘Optimization of ultrasoundassisted extraction of anthocyanins in red raspberries and identification of anthocyanins in extract using high-performance liquid chromatography-mass spectrometry’, Ultrasonics Sonochemistry, 14, 767–772. Cravotto G, Boffa L, Mantegna S, Perego P, Avogadro M, Cintas P (2008), ‘Improved extraction of vegetable oils under high-intensity ultrasound and/or microwaves’, Ultrasonics Sonochemistry, 15, 898–902. Csiktusnádi Kiss G, Forgács E, Cserháti T, Mota T, Morais H, Ramos A (2000), ‘Optimisation of microwave-assisted extraction of pigments from paprika powders’, Journal of Chromatography A, 889, 41–49. Da Porto C, Decorti D and Kikic I (2009), ‘Flavour compounds of Lavandula angustifolia L., to use in food manufacturing: comparison of three different extraction methods’, Food Chemistry, 112, 1072–1076. De Gennaro L, Cavella S, Romano R and Masi P (1999), ‘The use of ultrasound in food technology I: inactivation of peroxidase by thermosonication’, Journal of Food Engineering 39, 401–407. Dugand L R (1980), ‘Natural antioxidants’, In M.G. Simic, M. Karel (Eds), Autooxidation in food and biological systems, Plenum Press, New York, 261–295. El-Belghiti K and Vorobiev E (2004), ‘Mass transfer of sugar from beets enhanced by pulsed electric field’, Transactions of the Institution of Chemical Engineers, 82, 226–230. El-Belghiti K (2005), Effets d’un champ électrique pulsé sur le transfert de matière et sur les caractéristiques végétales, PhD thesis, UTC, Compiègne, France. El-Belghiti K, Moubarik R and Vorobiev E (2007), Use of moderate pulsed electric field, electrical discharges and ultrasonic irradiations to improve aqueous extraction of solutes from fennel (Foeniculum vulgare), unpublished data. El-Belghiti K and Vorobiev E (2005a), ‘Kinetic model of sugar diffusion from sugar beet tissue treated by pulsed electric field’, Journal of the Science of Food and Agriculture, 85, 213–218. El-Belghiti K and Vorobiev E (2005b), ‘Modelling of solute aqueous extraction from carrots subjected to a pulsed electric field pre-treatment’, Biosystems Engineering, 90(3), 289–294. Elkhori S, Jocelyn Paré J, Bélanger J, Pérez E (2007), ‘The microwave-assisted process: extraction and determination of fat from cocoa powder and cocoa nibs’, Journal of Food Engineering, 79, 1110–1114. El Zakhem H, Lanoisellé J-L, Lebovka N I, Nonus M and Vorobiev E (2006a), ‘Behavior of yeast cells in aqueous suspension affected by pulsed electric field’, Journal of Colloid and Interface Science, 300(2), 553–563. El Zakhem H, Lanoisellé J-L, Lebovka N I, Nonus M and Vorobiev E (2006b), ‘The early
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104 Separation, extraction and concentration processes stages of Saccharomyces cerevisiae yeast suspensions damage in moderate pulsed electric fields’, Colloids and Surfaces B: Biointerfaces, 479(2), 189–197. Eshtiaghi, MN and Knorr D, (2002), ‘High electric field pulse pretreatment: Potential for sugar beet processing’, Journal of Food Engineering, 52, 265–272. Fincan M, DeVito F and Dejmek P (2004), ‘Pulsed electric field treatment for solid–liquid extraction of red beetroot pigment’, Journal of Food Engineering, 64(3), 381–388. Fincan M and Dejmek P (2002), ‘In situ visualization of the effect of a pulsed electric field on plant tissue’, Journal of Food Engineering, 55, 223–230. Ganeva V, Galutzov B and Teissié J (2003), ‘High yield electroextraction of proteins from yeast by a flow process’, Analytical Biochemistry 315(1), 77–84. Grémy-Gros C, Lanoisellé J-L and Vorobiev E (2008) ‘Application of high-voltage electrical discharges for the aqueous extraction from oilseeds and other plants’, In: E. Vorobiev and N. Lebovka (Eds), Electrotechnologies for extraction from food plants and biomaterials, Springer, New York, 217–236. Grigonis D, Venskutonis P, Sivik B, Sandahl M, Eskilsson C (2005), ‘Comparison of different extraction techniques for isolation of antioxidants from sweet grass’, Journal of Supercritical Fluids, 33, 223–233. Grimi N, Lebovka M, Vorobiev E, Vaxelaire J (2009a), ‘Compressing behaviour and texture evaluation for potatoes pretreated by pulsed electric field’, Journal of Texture Studies, 40, 208–224. Grimi N, Lebovka M, Vorobiev E and Vaxelaire J (2009b), ‘Effect of a pulsed electric field treatment on expression behaviour and juice quality of Chardonnay grape’, Food Biophysics, 4, 1557–1858. Grimi N, Praporscic I, Lebovka N and Vorobiev E (2007), ‘Selective extraction from carrot slices by pressing and washing enhanced by pulsed electric fields’, Separation and Purification Technology, 58(2), 267–273. Grimi N, Vorobiev E and Vaxelaire J (2008), ‘Juice extraction from sugar beet slices by belt filter press: effect of pulsed electric field and operating parameters’, 14th World Congress of Food Science and Technology, 19–23 October 2008, Shanghai, China, TS24–29, 554. Gros C, Lanoisellé J-L and Vorobiev E (2003), ‘Towards an alternative extraction process for linseed oil’, Chemical Engineering Research and Design, 81(9), 1059–1065. Hemwimon S, Pavasant P and Shotipruk A (2007), Microwave-assisted extraction of antioxidative anthraquinones from roots of Morinda citrifolia, Separation and Purification Technology, 54, 44–50. Herrera M and Luque de Castro M D (2005), Ultrasound-assisted extraction of phenolic compounds from strawberries prior to liquid chromatographic separation and photodiode array ultraviolet detection, Journal of Chromatography A, 1100, 1–8. Hu A (2007), ‘Ultrasound assisted supercritical fluid extraction of oil and coixenolide from adlay seed’, Ultrasonics Sonochemistry, 14, 219–222. Imai T, Uemura K, Ishida N, Yoshizaki S, and Noguchi A (1995), ‘Ohmic heating of Japanese white radish Rhaphanus sativus L’, International Journal of Food Science and Technology, 30(4), 461–472. Jadhav D, Rekha B N, Gogate P R, Rathod V K (2009), ‘Extraction of vanillin from vanilla pods: A comparison study of conventional soxhlet and ultrasound-assisted extraction’, Journal of Food Engineering, 93, 421–426. Jaeger H, Balasa A, and Knorr D (2008), ‘Food industry applications for pulsed electric fields’, In: E Vorobiev and N Lebovka (Eds), Electrotechnologies for extraction from food plants and biomaterials, Springer, 181–216. Jemai A B and Vorobiev E (2002), ‘Effect of moderate electric field pulse (MEFP) on the diffusion coefficient of soluble substances from apple slices’, International Journal of Food Science and Technology, 37, 73–86. Jemai A B and Vorobiev E (2003), ‘Enhancing leaching from sugar beet cossettes by pulsed electric field’, Journal of Food Engineering, 59, 405–412.
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Principles of physically assisted extractions and applications 105 Jemai A B and Vorobiev E (2006), ‘Pulsed electric field assisted pressing of sugar beet slices: towards a novel process of cold juice extraction’, Biosystems Engineering, 93(1), 57–68. Kanakis C D, Daferera D J, Tarantilis P, Polissiou, M (2004), ‘Qualitative determintaion of volatile compounds and quantitative evaluation of safranal and 4-hydroxy-2,6,6,trimethyl-1-cyclohexane-1-carboxaldehyde (HTCC) in Greek saffron’, Journal of Agriculture and Food Chemistry, 52, 4515–4520. Kemp M R, and Fryer P J (2007), ‘Enhancement of diffusion through foods using alternating electric fields’, Innovative Food Science and Emerging Technologies, 8, 143–153. Kimbaris AC, Siatis N G, Daferera D J, Tarantilis P A, Pappas C S and Polissiou M G (2006), ‘Comparison of distillation and ultrasound-assisted extraction methods for the isolation of sensitive aroma compounds from garlic’, Ultrasonics Sonochemistry, 13, 54–59. Knorr D, Angersbach A, Eshtiaghi M N, Heinz V and Lee D-U (2001), ‘Processing concepts based on high intensity electric field pulses’, Trends in Food Science and Technology 12(3–4), 129–135. Knorr D, Geulen M, Grahl, T and Sitzmann W (1994), ‘Food application of high electric field pulses’, Trends in Food Science and Technology, 5, 71–75. Konduser M and Miklavcic D (2008), ‘Electroporation in biological cell and tissue: an overview’ In: E Vorobiev and N Lebovka (Eds), Electrotechnologies for extraction from food plants and biomaterials, Springer, New York, 1–38. Kulshrestha S and Sastry S K (2003), ‘Frequency and voltage effects on enhanced diffusion during moderate electric field (MEF) treatment’, Innovative Food Science and Emerging Technologies, 4, 189–194. Lagha A, Chemat S, Bartels P and Chemat F (1999), ‘Microwave–ultrasound combined reactor suitable for atmospheric sample preparation procedure of biological and chemical products’, Analusis, 27, 452–455. Lazarenko B R, Fursov S P, Scheglov Y A, Bordiyan V V, Chebanu V G (1977), Electroplasmolysis, Karta Moldavaneske, Kishinev, USSR (in Russian). Lebovka N I, Bazhal M I and Vorobiev E (2000), ‘Simulation and experimental investigation of food material breakage using pulsed electric field treatment’, Journal of Food Engineering 44, 213–223. Lebovka N I, Bazhal M I and Vorobiev E (2001), ‘Pulsed electric field breakage of cellular tissues: visualization of percolative properties’, Innovative Food Science and Emerging Technologies, 2, 113–125. Lebovka N I, Praporscic I and Vorobiev E (2004a), ‘Combined treatment of apples by pulsed electric fields and by heating at moderate temperature’, Journal of Food Engineering, 65, 211–217. Lebovka N I, Praporscic I and Vorobiev E (2004b), ‘Effect of moderate thermal and pulsed electric field treatments on textural properties of carrots, potatoes and apples’, Innovative Food Science and Emerging Technologies, 5, 9–16. Lebovka N I, Praporscic I, Ghnimi S and Vorobiev E (2005), ‘Does electroporation occur during the ohmic heating of food?’, Journal of Food Science, 70(5), E308–311. Lebovka N I, Shynkaryk M and Vorobiev E (2007), ‘Moderate electric fields treatment of sugarbeet tissues’, Biosystems Engineering, 96(1), 47–56. Li H, Pordesimo L and Weiss J (2004), High intensity ultrasound-assisted extraction of oil from soybeans, Food Research International, 37, 731–735. Li L, Ding L, Vorobiev E and Lanoiselle J-L (2009), ‘The application of high voltage electrical discharge for oil extraction from oil meals of linseed, rapeseed and palm kernel in aqueous solution’, Conference on Food Engineering CoFE 2009, April 5–8, 2009, Columbus, USA. Lianfu Z and Zelong L (2008), Optimization and comparison of ultrasound/microwave assisted extraction (UMAE) and ultrasonic assisted extraction (UAE) of lycopene from tomatoes, Ultrasonics Sonochemistry, 15, 2008, 731–737.
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106 Separation, extraction and concentration processes Lima M, Heskitt B and Sastry S (1999), ‘The effect of frequency and wave form on the electrical conductivity–temperature profiles of turnip tissue’, Journal of Food Process Engineering, 22(1), 41–54. Lima M, Heskitt B F and Sastry S K (2001), ‘Diffusion of beet dye during electrical and conventional heating at steady-state temperature’, Journal of Food Process Engineering, 24, 331–340. Lopez N, Puertolas E, Condon S, Raso J and Alvarez I (2009a), ‘Enhancement of the extraction of betanine from red beetroot by pulsed electric fields’, Journal of Food Engineering, 90(1), 60–66. Lopez N, Puertolas E, Hernández-Orte P, Alvarez I and Raso J (2009b), ‘Effect of a pulsed electric field treatment on the anthocyanins composition and other quality parameters of Cabernet Sauvignon freshly fermented model wines obtained after different maceration times’, LWT – Food Science and Technology, 42(7), 1225–1231. Luque-García J L and Luque de Castro M D (2004), ‘Ultrasound-assisted Soxhlet extraction: an expeditive approach for solid sample treatment: application to the extraction of total fat from oleaginous seeds’, Journal of Chromatography A, 1034, 237–241. Mandal V, Mohan Y and Hemalatha S (2008), ‘Microwave assisted extraction of curcumin by sample-solvent dual heating mechanism using Taguchi L9 orthogonal design’, Journal of Pharmaceutical and Biomedical Analysis, 46, 322–327. Mason T J (1990), Chemistry with ultrasound, Elsevier Applied Science, New York. McLellan M R, Kime R L and Lind K R (1991), ‘Electroplasmolysis and other treatments to improve apple juice yield’, Journal of Science of Food and Agriculture, 57(2), 303–306. Metaxas A C and Meredith R J (1993), Industrial microwave heating, IEEE, London. Palma M and Barroso C (2002), ‘Ultrasounds-assisted extraction and determination of tartaric and malic acids from grapes and winemaking by-products’, Analitica chimica acta, 458, 119–125. Pan X, Niu G, Liu H (2003), ‘Microwave-assisted extraction of tea polyphenols and tea caffeine from green tea leaves’, Chemical Engineering and Processing, 42, 129–133. Pan Y, Wang K, Huang S, Wang H, Mu X, He C, Ji X, Zhang J, Huang F (2008), ‘Antioxidant activity of microwave-assisted extract of longan peel’, Food Chemistry, 106, 1264–1270. Pérez-Serradilla J, Ortiz M, Sarabia L, Luque de Castro M (2007), ‘Focused microwaveassisted soxhlet extraction of acorn oil for determination of the fatty acid profile by GC-MS’, Analytical and Bioanalytical Chemistry, 388, 451–462. Povey M and Mason T J (1998), Ultrasound in Food Processing, Blackie Academic & Professional, London. Praporscic I, Ghnimi S and Vorobiev E (2005), ‘Enhancement of pressing of sugar beet cuts by combined pulsed electric field and ohmic heating’, Journal of Food Processing and Preservation, 29(5-6), 378–389. Praporscic I, Lebovka N I, Ghnimi S and Vorobiev E (2006), ‘Ohmically heated, enhanced expression of juice from apple and potato tissues’, Biosystems Engineering, 93(2), 199–204. Praporscic I, Lebovka NI, Vorobiev E, Mietlon-Peuchot M (2007), ‘Pulsed electric field enhanced expression and juice quality of white grapes’, Separation and Purification Technology, 52(3), 520–526. Praporscic I, Shynkaryk M, Lebovka N and Vorobiev E (2007b), ‘Analysis of juice colour and dry matter content during pulsed electric field enhanced expression of soft plant tissues’, Journal of Food Engineering, 79(2), 662–670. Priego-Capote F and Luque de Castro M D (2007), Ultrasound-assisted digestion: a useful alternative in sample preparation, Journal of Biochemical and Biophysical Methods, 70, 299–310.
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Principles of physically assisted extractions and applications 107 Raso J, Manas P, Pagan R and Sala FJ (1999), ‘Influence of different factors on the output power transferred into medium by ultrasound’, Ultrasonics Sonochemistry, 5, 157–162. Rogov I A and Gorbatov A V (1974), Physical methods of foods processing, Pischevaja Promyshlennost, Moscow (in Russian). Sastry S K (2005), ‘Advances in ohmic heating and moderate electric field (MEF) Processing’, In: G Barbosa-Canovas, M S Tapia and M P Cano (Eds), Novel food processing technologies, CRC Press, New York, 491–499. Schilling S, Alber T, Toepfl S, Neidhart S, Knorr D, Schieber A and Carle R (2007), ‘Effects of pulsed electric field treatment of apple mash on juice yield and quality attributes of apple juices’, Innovative Food Science and Emerging Technologies, 8, 127–134. Schreier P J R, Reid D G and Fryer P J (1993), ‘Enhanced diffusion during the electrical heating of foods’, International Journal of Food Science and Technology, 28, 249–260. Schrive L, Grasmick A, Moussière S, Sarrade S, (2006), ‘Pulsed electric field treatment of Saccharomyces cerevisiae suspensions: a mechanistic approach coupling energy transfer, mass transfer and hydrodynamics’, Biochemical Engineering Journal, 27, 212–224. Sensoy I and Sastry S K (2004), ‘Extraction using moderate electric fields’, Journal of Food Science: Food Engineering and Physical Properties, 69(1), 7–13. Sharma U, Singh H, Sinha A (2008), ‘Microwave-assisted efficient extraction of different parts of Hippophae rhamnoides for the comparitive evaluation of antioxidant activity and quantification of its phenolic constituents by reverse-phase high-performance liquid chromatography (RP-HPLC)’, Journal of Agriculture Food Chemistry, 56, 374–379. Shotipruk A, Kaufman P B and Wang H Y (2001), ‘Feasability study of repeated harvesting of menthol from biologically viable Mentha x piperata using ultrasonic extraction’, Biotechnology progress, 17, 924–928. Shynkaryk M V, Lebovka N I and Vorobiev E (2008), ‘Pulsed electric fields and temperature effects on drying and rehydration of red beetroots’, Drying Technology, 26(6), 695–704. Shynkaryk M V, Lebovka N I, Lanoisellé J-L, Nonus M, Bedel-Clotour C and Vorobiev E (2009), ‘Electrically-assisted extraction of bio-products using high pressure disruption of yeast cells (Saccharomyces cerevisiae)’, Journal of Food Engineering, 92( 2), 189–195. Sugiarto A T, Sato M and Skalny J D (2001) ‘Transient regime of pulsed breakdown in low-conductive water solutions’, Journal of Physics D: Applied Physics, 34(23), 3400–3406. Sun Y, Liao X, Wang Z, Hu X, Chen F (2007), ‘Optimization of microwave-assisted extraction of anthocyanins in red raspberries’, European Food Research Technology, 225, 511–523. Suslick K S (1988), ‘Ultrasound, its chemical, physical and biological effects’, VCH Publishers, New York. Van der Poel P W, Schiweck H and Schwartz T (1998), Sugar technology beet and cane sugar manufacture, Beet Sugar Development Foundation. Denver, USA. Virot M, Tomao V, Colnagui G, Visinoni F, Chemat F (2007), ‘New microwave-integrated Soxhlet extraction an advantageous tool for the extraction of lipids from food products’, Journal of Chromatography A, 1174, 138–144. Vorobiev E, Jemai A B, Bouzrara H, Lebovka N I and Bazhal M I (2005), ‘Pulsed electric field assisted extraction of juice from food plants’, In: G Barbosa-Canovas, M S. Tapia and M P Cano (Eds), Novel food processing technologies, CRC Press, New York, 105–130. Vorobiev E and Lebovka N I (2006), ‘Extraction of intercellular components by pulsed
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108 Separation, extraction and concentration processes electric fields’, In: J Raso and V Heinz (Eds), Pulsed electric field technology for the food industry. Fundamentals and applications, Springer, New York, 153–194. Vorobiev E and Lebovka N I (2008), ‘Pulsed electric field induced effects in plant tissues: fundamental aspects and perspectives of application’. In: E Vorobiev and N Lebovka (Eds), Electrotechnologies for extraction from food plants and biomaterials, Springer, New York, 39–82. Wang W C and Sastry S K (2002), ‘Effects of moderate electrothermal treatments on juice yield from cellular tissue’, Innovative Food Science and Emerging Technologies, 3, 371–377. Zeuthen P and Bogh-Sorensen L (eds) (2000), Food Preservation Techniques, CRC Press, New York. Zhang Q, Monsalve-Gonzalez A, Qin B L, Barbosa-Canovas G V and Swanson B G (1994), ‘Inactivation of Saccharomyces cerevisiae in apple juice by square wave and exponential-decay pulsed electric fields’, Journal of Food Process Engineering, 17, 469–478. Zimmermann U (1986), ‘Electrical breakdown, electropermeabilization and electrofusion’, Rev Physiol Biochem Pharmacol, 105, 175–256.
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Advances in process chromatography and applications 109
4 Advances in process chromatography and applications in the food, beverage and nutraceutical industries M. Ottens and S. Chilamkurthi, Delft University of Technology, The Netherlands Abstract: The basics of process chromatography and its application to separation and purification in the food, beverage and nutraceutical industry are discussed. Basic principles and modes of chromatography are described together with guidelines for appropriate modeling; all illustrated with examples. New developments are described in the area of novel ligands, new materials, and modes of operation and control. Finally, future trends with respect to the production of high-value nutraceuticals, regulation and process control are given. Key words: process chromatography, nutraceuticals, proteins, model, simulated moving bed.
4.1 Introduction 4.1.1 Chromatography as a unit operation in food, beverage and nutraceutical processing The global food, beverage, and agriculture industry caters to the population of the entire world. It is a complex, global collective of diverse businesses that together supply much of the food energy consumed by the world population. Global market forces are driving the continuous evolution of the food and beverage industry. Consolidation, changing consumer preferences and increasing government regulations are significantly impacting manufacturing and business strategy. Producing quality goods at the lowest possible cost is a major aim. Consumers are increasingly interested in the health benefits of foods and have begun to look beyond the basic nutritional benefits of food to the
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110 Separation, extraction and concentration processes disease-prevention and health-enhancing compounds contained in many foods. This, combined with a more widespread understanding of how diet affects disease, health-care costs and aging populations, has created a market for functional foods and natural health products. There is a growing interest in functional foods, which are dietary supplements that increase a person’s well being. A functional food contains one or more ingredients that are bioactive, that is: they promote desired biological activity in the human body. Process-scale chromatography is an indispensable unit operation in biotechnology and finds wide application in the food and beverage industry (Fig. 4.1). Even though the food industry does not rely on process scale chromatography to the same extent as perhaps other industries such as the pharmaceutical sector, the sheer size of the industry results in a considerable application for chromatography. Other than process-scale applications of chromatography, many applications are related to safety, quality control, and research of new food ingredients and products. Many product streams in the food industry contain components that may have applications for, for example, functional foods. These product streams are typically large in volume and contain only a small amount of the component (or mixture of components) of interest. For economical application of large-scale preparative protein, peptide or oligosaccharide chromatography from agricultural/biotechnology streams, there is a need for ‘cheap’ resins (the costs for resins may be as high as 70% of total processing costs). The standard technology for fractionation or enrichment of food streams is using a packed bed. The close contact between interstitial liquid and carrier beads ensures effective mass transfer. For the treatment of large process streams
Fermentation Cell disruption Separation of cell debris
Removal of ‘like’ molecules such as mis-folds, multimers, degradation products via: Crystallization (extraction) Chromatography
Concentration Purification
Physical adsorption Ion exchange Affinity, size (sorption processes)
Formulation
(membrane separations) (electrokinetic separations)
Market
Fig. 4.1 Positioning of chromatography in a general layout of a bioprocess. For food processes, the fermentation step may not be present.
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Advances in process chromatography and applications 111 and/or process streams where suspended solids or fouling components are present, problems may arise: the pressure drop over the column might be too high. The workhorse of preparative protein purification is ion exchange chromatography (IEC). For dilute feeds, IEC can be used in a capture step. Another way of pre-concentration is precipitation followed by resolubilization and IEC fractionation. Large-scale preparative chromatography can be done more efficiently using counter-current chromatography, which has larger driving forces for mass transfer and more efficient use of desorbent and (expensive) resin (see Section 4.4). As an example of the fractionation of protein hydrolyzates, simulated moving bed (SMB) chromatography has been recently applied (Ottens et al. 2006b). An example of large-scale purification of biomolecules being applied in industry is, for example, the purification of lysine (an amino acid) in SMB carousels (Van Walsum and Thomson, 1997). However, using a larger number of protein fractions, more side streams are needed, or a different approach will be needed (similar to distillation with different side streams). Resin screening is an important aspect of the development of an economical process (with respect to resin capacity, regeneration cycles, resin lifetime, costs of resins, food grade resins: some resin is always lost in the feed and the need for ‘water-soluble elution liquids’, as many food companies are not used to processing organic solvents). Last but not least, fouling characteristics are important as well as a proper cleaning program. The current challenges that are associated with the use of chromatography on a large scale in the food industry are the following (see also section 4.4): Large diluted stream with minor component (mg m–3) from large product streams (m3 h–1). – For high throughputs, expanded bed adsorption (EBA) or SMB chromatography might be considered, as indicated above. Considerable effort in the application and development of EBA was made by Kula and co-workers (e.g. Hubbuch et al., 2006). ∑ Pressure drop. Possible solutions include: – use of micro-structured materials, multimodal pores, perfusion resins; monoliths, structured woven packing, all allowing for a large convective flow and creating an interfacial area in the (functionalized) micropores, – (partly) fluidization, as in EBA, – radial flow chromatography. ∑ Fouling: – removal of debris, proteins and fats from the feed by membrane filtrations preferably in combination with a pretreatment. ∑ An understanding of the required purity of the targeted molecule. Mixtures with enriched fractions or pure targeted component. ∑
These considerations are addressed in this chapter.
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112 Separation, extraction and concentration processes Several examples of functional food products that can be enriched or purified by chromatography are briefly introduced next. Peptides are known to have a versatility of bioactivity, e.g. ranging from a metal binding function, through influencing mood functions, to a reduction of blood pressure. The peptides can be obtained from various (food) proteins; dairy liquids such as milk and whey have often been investigated (e.g. Meisel and Schlimme, 1996). Typically, the production of a bioactive ingredient for functional food is obtained by enzymatic hydrolysis of the native protein. From the resulting complex mixture, the peptide(s) of interest need to be isolated. Currently, this is often done by membrane filtration or chromatography. Membrane filtration is regarded as a cheap technology, but the selectivity is rather low. Electromembrane filtration has been suggested as an alternative isolation route for charged components. In this membrane-based separation, the selectivity is increased by applying an electrical field as the (selective) driving force for migration over the membrane (Bargeman et al., 2002a; 2002b), which contributes to separation costs and complexity in comparison to regular membrane filtration. On the other hand, chromatography is highly selective, but is often regarded as expensive, as a result of the relatively high costs of the column media and eluents. The use of SMB technology is investigated as an alternative separation route, as this counter-current chromatographic separation method has been shown to be able to efficiently reduce sorbent inventory and the use of buffers by one order of magnitude in comparison to regular (fixed bed) column chromatography (e.g. Ballanec and Hottier, 1993; Cavoy et al., 1997). Furthermore, the use of chromatography for glucose–fructose separation (Azevedo and Rodrigues, 2005; Luz et al., 2008) sugar recovery from molasses (Hannu and Jarmo, 2000), stabilization of beers (Rehmanji et al., 2002), separation of proteins from various sources, such as milk (Kim et al., 2003a) and rapeseed, are only a few of the process scale applications of chromatography in the food and beverage industry. In addition to this several other functional components such as amino acids (Kostova and Bart, 2007a,b), flavors and fragrances (Gunther and Armin, 1993), nutraceuticals (Genovese et al., 2007), organic acids (Yoshikawa, et al., 2007), and oligosaccharides (Berensmeier and Buchholz, 2004) have also been separated on process and analytical scales. All these separations involve various interaction mechanisms, including ligand-exchange chromatography, LEC (Azevedo and Rodrigues, 2005, Luz et al., 2008, Steffenson and Westerlund, 1996), size-exclusion chromatography, SEC (Thurl et al., 1993), IEC (Kim et al., 2003a), and reversed-phase highperformance liquid chromatography, RP-HPLC (Stah et al., 1994; Thurl et al., 1991). The various interaction modes are described in the next section. Sanitation and hygienic design are key characteristics of food, beverage and nutraceutical processing. Major chromatographic companies that operate in the field are amongst others Universal Oil Products (UOP) and Novasep, and they can readily assist in designing large-scale chromatographic separations.
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Advances in process chromatography and applications 113 4.1.2 Main chapter themes In the remaining part of this chapter first the basic principles of chromatography are described and the various types and basic modeling outlined. Then some applications are presented to indicate how chromatography is currently employed and its future potential in the food field. Subsequently, a section on recent developments in the chromatographic field is given with implications and applications for use in the food industry. This section is followed by an overview of future trends in the food chromatography field and finally some conclusions and suggestions for further reading are given.
4.2 Basic principles of process chromatography 4.2.1 Types of chromatography used The different modes or types of chromatography that can be used are categorized and depicted in Fig. 4.2 (see also GE website for further information; GE Healthcare is one of the major suppliers of chromatography equipment and resins: http:\\www.gehealthcare.com). The various methods are ranked according to their interaction strength, with the weakest on top. Size-exclusion chromatography is based on differences in the size of the solute molecules, by using a porous resin with distinct pore sizes (Horneman et al., 2004, 2006). Hydrophobic interaction chromatography and reversed-phase chromatography involve separation based on the differences in (surface) hydrophobicity of the solutes. The most powerful and most often used in food processing is ion-exchange chromatography, which separates solutes based on the differences in charge, which can be modulated by the pH and ionic strength of the solution. The strongest and most selective method is Sorption mechanisms
Weak
∑ Size
∑ Gel filtration (GF)
∑ Van der Waals
∑ Hydrophobic interaction (HIC) and reversed phase (RPC)
∑ Polar
d+
d–
d+ d+ +
∑ Ionic Strong ∑ Affinity
+
+
∑ Ion exchange (IEC)
+ ∑ Affinity chromatography (AC)
Fig. 4.2 Illustration of the basic interaction modes of chromatographic separation.
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114 Separation, extraction and concentration processes affinity chromatography, and it is most often used in the pharmaceutical field, because of its selectivity and because its high cost can be met here. The purpose of process chromatography is to separate a target component from a mixture. Liquid chromatographic separations are based upon the different degrees of interaction of solutes dissolved in a mobile phase with a chromatographic medium or resin (the stationary phase), resulting in differential migration of the solutes through the column as illustrated in Fig. 4.3. The feed to the chromatography column is a binary mixture of solutes A and B, moving with concentration (wave) velocities wA and wB, respectively, through the column, with wB > wA. That is, solute B is more strongly retained or retarded by the resin than solute A. Each solute band becomes more dispersed as the chromatographic process progresses (depicted with fainter shading in Fig. 4.3b). Therefore, a chromatographic purification that is mainly based on adsorption and desorption of the target compound consists of several steps as follows. Equilibration of the resin For equilibration of the resin, the adsorbent material is prepared for the chromatographic separation. Depending on the mechanism of adsorption the equilibration step may include equilibration with the desired ionic strength, pH, and solvent concentration. Usually the equilibration solvent is chosen such that upon adsorption of the target compound and its impurities no change in composition (pH, ionic strength, and solvent composition) of the mobile phase occurs. In some instances it may be necessary to change the composition of the mobile phase to enhance the adsorption process. Loading In the loading process step, the target compound (and some impurities) is adsorbed onto the media. The capacity of the resin and/or the difference in affinity between target compound and its impurities determine the loading volume and loading conditions. It may be necessary to change or alter the feed stream properties to optimize the adsorption process. Washing In the wash phase, a washing buffer or solvent is used to desorb impurities from the resin. The washing buffer may be similar to the equilibration buffer, but this is not necessarily the case. Elution of the target compound needs to be avoided in this step. Elution In the elution phase the target compound is desorbed into an elution buffer solution. This can be done both in isocratic mode (constant mobile phase composition) and in gradient mode in which a gradient in the eluting buffer is applied to achieve optimal purity and/or minimal elution volume.
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Advances in process chromatography and applications 115 Feed Mobile liquid phase
WB
Mobile phase
WA
WB
Mobile phase
WA
WA
WB
More strongly retained Less strongly retained
Stationary solid phase
t = 0 min
t = 10 min
t = 20 min Concentration
tR,B-tR,A
A
(a)
B A
B A
tR,A
tR,B
wA
wB
Time
Regenerant (eluent)
Feed B A
B
Flush
A
A B
B
B
A
A
B
A
A
A
A
Time
(b)
Fig. 4.3 (a) Illustration of the basic principle of chromatographic separation. Chromatogram showing the exit concentrations of A and B as a function of time (arbitrary times are used to show the progress of the chromatographic separation). (b) Adsorption, flush and regeneration.
Sanitization After elution a sanitization step may be applied in which compounds that show a higher affinity to the resin than the target compound are desorbed. For polymeric media, often a solvent is used to remove very apolar compounds from the resin. For hydrophilic media, a caustic solution is often used to © Woodhead Publishing Limited, 2010
116 Separation, extraction and concentration processes hydrolyze any irreversibly bound proteins. Acidic washes are applied to remove salts with low solubility (often calcium salts). Varying the salt concentration, pH or solvent composition can also be used to deliberately shrink and swell the resin which may enhance sanitization. After sanitization, the process starts again from equilibration. For processes in which there is very little affinity difference between the target compound and a limited number of impurities, it is often decided to use a continuous mobile phase with constant composition. The compounds are collected sequentially at the exit of the column and no intermediate equilibration, washing and sanitization is applied. A typical example of such a process is the separation of glucose and fructose on strong cationexchange resins in the Ca form. The requirements for sorbents to be used in chromatography and the different basic types are outlined in Fig. 4.4. 4.2.2 Modes of operation There are three different modes of operating a chromatographic process: frontal, elution and displacement. These stem from the three different ways of desorbing bound solutes from the chromatographic medium. Requirements ∑ Insoluble
∑ Hydrophilic Compromise
∑ Macroporous ∑ Mechanically stable
∑ Proper shape ∑ (Bio-)chemical stability
Main types
Composite
Organic ‘polymers’
Inorganic ‘pellets’
(a) ∑ Inorganic
∑ Organic polymers
Activated carbon
Polystyrene
Silica
Poly(meth)acryl
Alumina
Polysaccharides (dextran, agarose, etc.)
Pellet
Dense gel
Apolar
Macroreticular
Polar
Open gel
(b)
Fig. 4.4 (a) Requirements for sorbents and the main types available. (b) Examples of inorganic and organic polymers used as sorbents.
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Advances in process chromatography and applications 117 In frontal chromatography, the sample is also used for elution. In this case, the sample feed is continuously applied to the column and the sample components displace each other in order of decreasing affinity for the chromatographic medium. The least retained solute is obtained in a pure form (i.e. depleted of the more strongly retained components) until the other solutes break through. Eventually, the column is saturated with the strongest retained component and the effluent has the same composition as the feed. In elution chromatography, the solutes are desorbed from the chromatographic medium by the action of a competing or modifying agent in the mobile phase, such as a salt. The concentration of the modifying agent is a key parameter. Three situations can be distinguished: isocratic, gradient or step change in the mobile phase composition. By keeping the composition (e.g. ionic strength) of the eluting buffer constant, the capacity factor is also kept constant; this is called isocratic elution and it is primarily used for separation of small molecules for which the variation of retention factor with mobile phase composition is not as large as for larger molecules. In gradient elution, the composition of the mobile phase is continuously changed (e.g. increasing ionic strength as in ion exchange chromatography), resulting in a continuous change in the retention factor. Instead of using a continuous gradient, the mobile phase composition may be changed in a stepwise fashion, resulting in discontinuous changes or ‘shifts’ in the retention factor. Gradient or step elution may be necessary for separation of large biomolecules for which the binding strengths may be too strong for isocratic elution to be feasible. In displacement chromatography, the affinity of the competing agent for the chromatography media is much higher than that of the solute, such that the agent effectively displaces the solute. This takes place irrespective of the concentration of the displacing agent (or displacer), in contrast to elution chromatography. Displacement chromatography has the advantage of being able to concentrate samples. However, selecting an appropriate displacer may not be easy. Several important aspects of using chromatography in food separations are outlined in Fig. 4.5. 4.2.3 Basic modeling The use of mathematical modeling is helpful in the design of chromatographic separators to size the chromatographic piece of equipment and to determine its optimal operation conditions. There is a wealth of information available on modeling of chromatography in different levels of detail. Guiochon gives a good basic overview (see section 4.8), and some recent modeling efforts using the general rate model and the linear driving force model for mass transfer are given by Nilsson (2005) and Horneman et al. (2006), respectively. Although the general rate model is the more accurate one, a less computationally involved but still accurate one is the linear driving force
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118 Separation, extraction and concentration processes Eluent recycle ∑ Operating conditions and short-cut design Raffinate product Feed
∑ Complex phase behavior (azeotropic separations) ∑ Non-equilibrium fractionation ∑ Reactive separations
Extract product
Eluent make-up
∑ Multiple side/product streams ∑ Column dynamics
Sorbent recycle
Fig. 4.5 Schematic diagram of the operation of a large-scale chromatographic fractionator, such as a simulated moving bed, and the main factors for successful operation.
model, as shown in Fig. 4.6. It contains two partial differential equations describing the concentration of solutes in the bulk liquid and solid phase as function of time and axial position in the column. The equations are connected by a mass transfer term containing the isotherm, which gives the ratio of partitioning over the solid and liquid phase of a solute under certain environmental conditions. Several isotherms can be used, the most common being the Langmuir one, which gives a linear relation between the concentration in the solid and liquid phase at low concentrations and levels off at a constant saturated value at higher concentration at the maximum loading capacity of the specific resin used. Any appropriate PDE solver can be used (e.g. PDESol, Comsol) to solve the set of obtained partial differential equations. Commercial packages that are available on the market are for example ASPEN Chrom, gPROMS, and for flowsheeting SuperProDesign (Intelligen).
4.3 Applications of process chromatography in the food, beverage and nutraceutical industries Process scale chromatography has been effectively used in the food, beverage and nutraceutical industry (Fig. 4.7 and Table 4.1). Examples of these applications are presented here and by doing so a brief overview is provided of the industrial areas where process chromatography is applied. The various sources and products that can be concentrated, fractionated and purified by chromatography are given in Fig. 4.7 and Table 4.1, as well as scales of some chromatographic separations actually in industrial practice with its operation mode (SMB, and fixed bed chromatography, FB).
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Advances in process chromatography and applications 119 4.3.1 Glucose–fructose separation The technology used previously (up to the middle 1970s) in the preparation of fructose was based on the application of a scarce and expensive raw material, inulin, which is a fructose polymer. This prevented the wide-scale use of fructose in the food and medical industry. Only in the 1980s did the application of chromatographic methods permit the large-scale manufacture of fructose by its isolation from glucose–fructose syrups obtained either as ∑ ∑ ∑ ∑
Which phases? Which convective streams? Is there dispersion? Mass transfer between phases?
Void fraction f
(Interstitial) velocity v
Feed flow Q MT
Cross-section A
Q = v fA
Mass transfer
Dispersion
(a) Mass balance over a slice Convection
(fA)vc
Dispersion
Ê ˆ (fA)Á – E ∂c ˜ Ë ∂z ¯
Mass transfer Accumulation
z + Dz
– [(1 – f)A] MTDz (fADz ) ∂c ∂t
Change = in – out ∂c = – ∂t Equation 1
z
Per volume solid
Dz (vc ) |z+ z
Dz
+E
Ê ∂c ˆ ÁË ∂z ˜¯
∂c = – v ∂c + E ∂2c – 1 – f MT ∂t ∂z f ∂z 2
z+Dz z
Dz
–
1–f MT f
E = 2vdp
(b)
Fig. 4.6 Mathematical modeling for chromatography (a). Different (partial differential) equations to be used (b) and (c), with ‘closing relations’ for mass transfer (d). Symbols: z axial length scale, D difference, c concentration, t time, k mass transfer coefficient, a interfacial area, K partition coefficient, Sh Sherwood number, Re Reynolds number, Sc Schmidt number, D diffusion coefficient, d particle diameter, n kinematic viscosity, v velocity, A cross- sectional area, Q flow rate, f void fraction, MT mass transfer, E dispersion coefficient. Subscripts: f film, p particle. Superscripts: * at equilibrium, vinculum solid phase.
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120 Separation, extraction and concentration processes z + Dz
z
Mass balance over a slice of solid Mass transfer MTDz = ka (c* – c) Dz Overall coefficient Area per volume of solids Equilibrium concentration Accumulation
∂c Dz ∂t
Equation 2
∂c = ka (c* – c) = MT ∂t
Overall MT coefficient k–1 = (Kkf)–1 + k–1 p
Solve two coupled PDEs (c) dp = 10–3 m
In the ‘film’ Shf = 4 + 1.1Re0.6 Scf0.33 kf dp
Shf ∫ Re ∫
Df
; Scf ∫
nf ; Df
vdp nf
Df = 10–9 m2 s–1 v = 10–3 m s–1 Re = 0.5 Scf = 1000 Shf = 10 kf = 10
Df dp
= 10–5 m s–1
In the ‘particle’ Dp = 10–10 m2 s–1
Shp = 10 Shp ∫
k pd p Dp
; Dp π Df
kp = 10
Dp dp
= 10–6 m s–1
(d)
Fig. 4.6 Continued
a result of the inversion of sucrose or by converting glucose with the aid of glucose isomerase. The most common type of fructose syrup, usually called high-fructose syrup, contains 42% fructose, 52% glucose, and 6% oligosaccharides on a dry basis, and is used in place of sucrose in some foods and beverages (Kishihara et al., 1992). For some purposes, because fructose is sweeter and more soluble in water at low temperatures than glucose, syrup with 55 to 90% fructose, called higher-fructose syrup, is desirable. Therefore, to produce syrup containing more than 50% fructose a process to separate fructose from the equilibrium mixture is desirable. This mixture can be separated by chemical and physical means. The most widely used method for this kind of separation is chromatographic separation using cation exchange resins in Ca2+ form (Barker et al., 1984). In this instance, SMB is a widely used technology (see Section 4.4.3) (Bubnik et al., 2002). © Woodhead Publishing Limited, 2010
Advances in process chromatography and applications 121
Dairy industry
Marine
Milk Eggs
Fish Algae Peptides Proteins
Omega-3, -6 fatty acids
Agriculture Sugars Flavors
Fig. 4.7 Source and product areas where process chromatography is being used or has potential. Table 4.1 Industrial scale chromatographic separations. Taken from Lecture notes from the Advanced Course on Downstream Processing, Delft, 2009, and Ganetsos and Barker (1992) (see Section 4.8) System
Process*
Scale (tonne a–1)
Xylene isomers Glucose–fructose Lactic acid, citric acid Amino acids Chiral compounds Peptides Milk proteins Flavors
SMB SMB FB/SMB FB/SMB SMB SMB FB FB
400 000 100 000 10 000 1500 1–10 0.1–1 0.1 0.01–01
*FB, fixed bed; SMB, simulated moving bed.
When large-scale separations are required, the SMB process is often the technology of choice in order to reduce separation costs. Compared with batch chromatography, SMB technology uses less eluent; in addition the separation productivity is higher thus using less adsorbent (Broughton, 1984; McCulloch et al., 1994). In recent years, fructose and glucose were separated from complex mixtures such as cashew apple juice using SMB technology (Azevedo and Rodrigues, 2005; Luz et al., 2008). 4.3.2 Recovery of sucrose from molasses Chromatographic separations are used to separate valuable components from beet molasses in beverage industries. Characteristic products that are © Woodhead Publishing Limited, 2010
122 Separation, extraction and concentration processes currently produced in addition to sugar include betaine, inositol, peptides, amino acid mixtures and several individual amino acids. The recovery of the other components can significantly improve the economy of sucrose recovery, but still sucrose remains as the most important product controlling the overall economy of the process. Chromatographic separation can also be applied to treat intermediate juices in the beet sugar industry. Thick juice separation has also been suggested. The resolution of the peaks may change owing to the changing composition of the feed material. Industrial batch systems were first built in the 1960s and the 1970s. In the mid-1980s the continuous simulated moving bed process was applied in molasses separation. The sequential SMB process, which can simultaneously recover multiple product fractions, was introduced in the late 1990s (Hyöky, 1999; Lancrenon, 1997; Paananen, 1996; Rousset, 1997). Hyöky et al. (1998) described the multi-profile FAST separation, which effectively doubled the efficiency. Further development includes a patented two-stage process (Hyöky et al., 1998). It is a combination of two chromatographic fractionators to improve the recovery and the purity of overlapping components, such as sucrose and betaine. It is uniquely suitable for recovering multiple value-added products from the process streams of the sugar industry (Paananen and Kuisma, 1999). Chromatographic systems are used in production of sweeteners based on beet, cane and starch as described by Paillat (1999). Sugar beet produces sucrose, betaine, raffinose, invert and many other water-soluble components. These components are extracted from the beets as juice. Part of these watersoluble components are removed or destroyed in the juice purification, but a large part of them will end up in the by-product: molasses. Molasses is the most common raw material in chromatographic desugarization plants. The main product is sucrose, but in addition betaine, inositol, amino acid mixtures and individual amino acids have been commercially separated from molasses (Hyöky et al., 1998; Paananen, 1996; Paananen and Kuisma, 1999). Low green or B molasses separation is already in industrial use. Kearney has suggested thick juice separation (Kearney, 1997; Kearney and Rearick, 1995). The positioning of the chromatographic step in the beet process has distinct effects on the recovery of sucrose and especially on the recovery of the other components. 4.3.3 Stabilization of beers Brewery production of some corns and hops for beer is a universal process in the beer industry. The quality of beer depends on its taste, clearness, color and foam retention. One of the key processes is the procedure of removing non-micro-organisms impurities. Fixed-bed chromatography with several resins like silica gel (SiO2nH2O) are in use to remove proteins that make beer muddy. Beers are known to contain a wide variety of phenolic compounds, most of which originate from the raw materials of brewing, i.e., barley and hops
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Advances in process chromatography and applications 123 (Gardner and McGuinness, 1977). Of these phenolic compounds, the flavanoids are of particular interest to brewers, as they have long been thought to be precursors of non-biological haze in unstabilized beers (McMurrough et al., 1983). The flavanols have been classified into three groups on the basis of their chromatographic behavior. The first group, the simple flavanols, comprise flavanol monomers [e.g. (+)-catechin and (–)-epicatechin)], dimers (e.g. prodelphinidin B3 and procyanidin B3), and trimers. The second groups, the polymeric flavanols, are formed by oxidation and polymerization of simple flavanols. The complexed flavanols, the third group, result from the interaction of polyphenols with proteins to form complex structures (McMurrough, 1979). Stabilization of beer against haze formation may be achieved by decreasing the simple flavanol content, thereby limiting further flavanol polymerization and complexation. Increasingly, this stabilization is being achieved in breweries by treatment of beer with polyvinylpolypyrrolidone (PVPP) before packaging (McMurrough, 1993; McMurrough et al., 1992). Stabilization is achieved by sorption of phenolics on PVPP, which is subsequently removed by sheet filtration. This procedure can be carried out in conjunction with silica hydrogel treatment to remove haze-forming proteins (McMurrough and Madigan, 1996). These amorphous, non-additive stabilizers offer selective removal of haze forming proteins, improving colloidal stability of beer with no adverse effects on head or taste. Typically these silicas have been used with body feed diatomaceous earth filter aid (D.E.). However, by changing the conditions in the manufacturing process, it is possible to vary permeability (filterability) of silica hydrogel products. An innovation in the stabilization of beer is the ‘combined stabilization system’ (CSS) which is capable of a combined removal of turbidity forming protein and polyphenols in a single step (Janey and Katzke, 2002). As the economic alternative to PVPP and silica stabilization, the CSS is a compact rig mounted fully automated stabilization system, which can be integrated into any existing filter line. The adsorbent is permanently retained between an inlet and an outlet screen, making the dosing of precoat and feed suspension before each filtration unnecessary. A CSS adsorber has a particle size of 100– 300 mm. The solid phase adsorber resin is based on a high-grade, cross-linked, insoluble agarose (polysaccharide). Protein and polyphenols are adsorbed and then removed from the agarose adsorbent during the regeneration. Neither substance is dissolved in the beer nor is the beer quality affected in a negative way. The beer’s organoleptic properties, its foam stability, color and bitterness units remain unaltered with this kind of approach. 4.3.4 Separation of lysozyme from egg white Lysozyme has a number of possible applications; for example it can be used as an additive to baby milk or ophthalmic preparations; for treatment of ulcers, wounds and infections, as a potentiator of some antibiotics; as an
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124 Separation, extraction and concentration processes antioxidant; and for its antimicrobial properties (Owen and Chase, 1997). The expanding potential for application of lysozyme in many fields of science and technology dictates the development of efficient and simple methods for lysozyme purification. Several chromatographic techniques that have been used for the separation of hen egg white lysozyme include ion exchange chromatography (Banka et al., 1993; Li et al., 1999; McCreath et al., 1997), affinity chromatography (Chiang et al., 1993; Yamada et al., 1985; Yamasaki and Eto, 1981), dye-binding chromatography (Tejeda-Mansir et al., 2003), affinity membrane separation (Arica and Bayramoglu, 2005; Bayramoglu et al., 2003; Ruckenstein and Zeng, 1997), ultrafiltration (Ghosh and Cui, 2000; Ghosh et al., 2000), PEG/salt aqueous two-phase system (Su and Chiang, 2006), reverse micelles (Noh and Imm, 2005), metal-affinity precipitation (Roy et al., 2003) or adsorption to plant residues (Hou and Lin, 1997). Commercial cation-exchange resins such as Duolite C-464, Amberlite, CM Sephadex have been used for the separation of lysozyme on process scales with recoveries of 90–95% (Li-Chan et al., 2006). Nonetheless, the lengthy steps and dilution of product at the end of the process have hindered their application at process scales. More recently this problem has been overcome by the use of magnetic separation techniques for the separation of egg white lysozyme and several other proteins (Safarik and Safarikova, 1993; 2004). These techniques usually enable simple, onestep separation of target proteins, and in most instances magnetic affinity adsorbents can be used. However, large-scale application may be difficult to achieve. 4.3.5 Separation of whey protein from milk Milk proteins are the most important source of bioactive peptides. Typical bovine milk contains 13% solids, with 4% fat present as an emulsion of globules with diameters up to 10 mm and caseins present as a colloidal suspension of particles with diameters up to 0.1 mm (Bylund, 2003). Fat globules normally cause problems for chromatographic separations, as they block packed columns as soon as the feed is introduced. Raw whole milk contains larger suspended particles than whole (full fat) processed milk, as the latter is homogenized to produce a uniform consistency. The composition of bovine milk includes water, fat, lactose and minerals, and up to 6% of the mass is made up by proteins and peptides, among them a number of high-value substances (Lourdes et al., 2000). In particular, milk contains two major protein groups, caseins and whey proteins, which differ greatly with regard to their physicochemical and biological properties. Normal milk contains 30–35 g L–1 proteins, approximately 80% of which are caseins with the remainder being the whey proteins (Korhonen et al., 1998). Whey proteins can be acquired as a by-product in cheese manufacturing process. The required long-term stability in functional performance of these proteins is often lacking. In general, whey is dilute liquid composed of lactose, a variety
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Advances in process chromatography and applications 125 of proteins, minerals, vitamins and fat. Whey contains about 6% solids of which 70% or more is lactose and about 0.7% is proteins (Gerberding and Byers, 1998). Whey protein components are a-lactalbumin, b-lactoglobulin, immunoglobulins A, M and G, bovine serum albumin (BSA), lactoferrin and lactoperoxidase. b-Lactoglobulin is the major whey protein in bovine milk; it has a molecular weight of 18.4 kDa, possesses 162 amino acid residues and its concentration is 2–4 g L–1. a-Lactalbumin is an albumin which has 123 amino acid residues. It possesses a molecular weight of 14.2 kDa and its concentration in milk is 0.6–1.7 g L–1 (Ye et al., 2000). Bovine whey proteins have potential applications in veterinary medicine, food industry and as supplements for cell culture media. Immunoglobulin G (IgG) and immunoglobulin A (IgA), present in bovine whey, have high pharmaceutical value (Hahn et al., 1998). a-Lactalbumin can be used in infant formula and as a nutraceutical because of its high tryptophan content. b-Lactoglobulin is used in the production of confections (Zydney, 1998). Oral administration of bovine IgG is known to be an effective treatment of various infections of newborn infants (Hutchens et al., 1990). Typical concentrations, molecular weights and isoelectric points of whey proteins are given in Table 4.2 (Al-Mashikhi et al., 1988). Several techniques are used for the partitioning of whey proteins by e.g. aqueous two-phase systems (Jose et al., 2000; da Silva and Meirelles, 2000; Rito-Palomares and Miguel, 1998). Also, some attempts have been made to isolate whey proteins by using membrane filtration (Lucas et al., 1998; Gerd et al., 2000). The separation of whey proteins by using ion exchange chromatography on a process scale has been investigated by many researchers and several methods have been reported (Gerberding and Byers, 1998); Ionexchange separations take advantage of electrostatic interaction between surface charges on biomolecules, such as amino acids or proteins, and clusters of charged groups on the resin phase. An adsorbing biomolecule displaces counterions associated with the surface, discharging a complementary buffer salt in the process. Adequate buffering is required to shield native protein structures from changes in pH adjacent to exchange surfaces (Donnan effect) and pH effects were induced by sorption. Selection of an appropriate buffer Table 4.2 Typical concentration of whey proteins and their isoelectric points (pI) Protein
Approximate MW (kD) concentration (w/w%)
Isoelectric point (–)
b-Lactoglobulin a-Lactalbumin Immunoglobulins (A, M, G) BSA Protease-peptones Lactoferrin (LF) Lactoperoxidase (LP)
0.3 0.07 0.06 0.03 0.14 0.003 0.002
5.35–5.49 4.2–4.5 5.5–8.3 5.13 3.3–3.7 7.8–8.0 9.2–9.9
18.4 14.2 150–900 69 37–55 78 78
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126 Separation, extraction and concentration processes is critical to the success of ion-exchange separations (Keith and Edwin, 1995). 4.3.6 Separation of lactoferrin and lactoperoxidase In addition to the caseins and whey proteins, which are generally isolated from milk, some specialty proteins are also identified as byproducts in these separations. Two very important such proteins of commercial interest, which are part of the whey protein fraction, are lactoferrin (LF) and lactoperoxidase (LP). Although production of such high-value whey proteins is a commercial reality, two aspects of dairy processing may not be optimal for their production. First, the proteins are subjected to a series of processing steps before being extracted. It is a generally accepted principle of bioseparation process design that proteins should be separated from a source material as fast and in as few steps as possible to avoid loss of activity and yield (Ladisch, 2001; Harrison et al., 2003; Ottens et al., 2006a). Lactoferrin (LF) and lactoperoxidase (LP) being basic proteins are usually captured from whey or skim milk by cationexchange chromatography and sold as specialty ingredients (Lonnerdal and Carlsson, 1977). The costs associated with the separation of these types of proteins is high, but the value and wide range applications in various functional food and nutraceutical products outweigh the cost. Lactoferrin, a 70–80 kDa single subunit glycoprotein, consist of two domains, each with one iron-binding site that requires synergistic binding of a bicarbonate anion (Al-Mashikhi and Li-Chan, 1988). A conformational change in the binding site, between open or closed forms, accompanies iron binding and release. The affinity of lactoferrin for iron is much greater than of transferrin which functions as an iron carrier in serum. It is present in human milk, and many exocrine fluids (saliva, mucous secretion) as well as in mammalian. Lactoferrin is a complex molecule with a number of interesting functional properties. Based upon these, lactoferrin may enter the composition in a wide range of products. Lactoperoxidase belongs to the category of peroxidase and acts as a biopreservative. Both these products find several applications based on their antimicrobial and antiviral activities such as in baby foods, food supplements, cosmetics, oral care, and preservation of meat. The production of high-value dairy proteins such as lactoferrin and lactoperoxidase normally requires extensive pre-treatments of milk to remove fat and caseins by centrifugation, precipitation, Ca2+ chelation and/ or filtration. Similarly, fat and caseins are normally removed before capture of recombinant proteins from the milk of transgenic animals (Morison and Joyce, 2005). Such pre-treatments can result in significant loss of protein yield and/or activity. Currently, high-value dairy proteins are viewed as by-products, with the major income of the industry coming from commodity dairy foods such as milk powder, cheese and butter. Economies of scale for production of commodity
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Advances in process chromatography and applications 127 dairy products mean that centralized processing is the industry norm. Milk is typically cooled and held at 4 °C in vessels on the farm for up to 2 days before being transported to a dairy factory. There, it is pumped to holding tanks and then undergoes a series of unit operations such as cream (fat) separation, pasteurization, homogenization and blending for standardization before further processing into individual products (Bylund, 2003). After a number of such operations (which varies from factory to factory), LF and LP may be extracted from skim (low fat) milk or, more commonly, from whey, which is produced as permeate during membrane concentration of milk or after precipitation of caseins much further down the processing chain as a by-product of casein production or cheese making (Tomita et al., 2002). Extensive pre-treatments of milk and whey before ion-exchange capture of proteins are by no means restricted to industry but are also used in the laboratory. Many authors have examined the capture and analysis of whey proteins by chromatography (Al-Mashikhi and Li-Chan, 1988; Al-Mashikhi and Nakai, 1987; Andrews et al., 1985; Chaplin, 1986; de Frutos et al., 1992; Donnelly, 1991; Doultani et al., 2004; Elgar et al., 2000; Etzel et al., 2000; Felipe and Law, 1997; Francis and Regester, 1995; Geberding and Byers, 1998; Hahn et al., 1998; Humphrey and Newsome, 1984; Konecny et al., 1994; Lonnerdal and Carlsson, 1997; Morr and Ha, 1993; Noppe et al., 1999; Torre and Cohen, 1996; Visser et al., 1991; Xu et al., 2000; Ye et al., 2000; Yoshida, 1988). Hahn et al. (1998) examined the performance of a range of commercially available pharmaceutical grade cation exchangers for protein capture from acid whey. Doultani et al. (2004) used cationexchange chromatography to produce a number of protein products from mozzarella cheese whey. Ye et al. (2000) used both anion and cation exchange chromatography to isolate a-lactalbumin, b-lactoglobulin, lactoferrin and lactoperoxidase from rennet whey. Cation exchange membrane chromatographic systems have been successfully used for the separation of lactoferrin and lactoperoxidase on an industrial scale from sweet whey concentrate with yields of more than 90% for lactoferrin. It has also been demonstrated that these system can be scaled up to handle 1 ¥ 105 tonne/year of whey (Plate et al., 2006). 4.3.7 Separation of napin from rapeseed meal Preparative chromatography of napin and other proteins from rapeseeds was successfully carried out with a combination of ion exchange followed by hydrophobic interaction and size-exclusion chromatography. The yield of this combined process was 18% for napin and 40% for cruciferin (Berot et al., 2005). Rapeseed proteins were underexploited as it was being used only as animal feed until its non-food applications were demonstrated in a European project (Green chemicals and biopolymers from rapeseed meal with enhanced end-user performance. European contract Enhance QLK5 CT 199901442). Hence large quantities were needed to demonstrate its potential © Woodhead Publishing Limited, 2010
128 Separation, extraction and concentration processes in functional additives. Preparative purification processes of seed storage proteins were demonstrated for 11S and 7S globulin-type proteins in pea and soy (Crevieu et al., 1996; Gueguen et al., 1984; Howard et al., 1983; Lehnhardt et al., 1983; Larre and Gueguen, 1986; Nielsen, 1985; Thanh and Shibasaki, 1976). But with rapeseed storage proteins, which are composed of 11S globulin and 2S albumin-type proteins, only analytical purifications were developed, except procedures based on ammonium sulfate selective precipitation of proteins (Raab and Schwenke, 1984). SEC was successfully used to achieve the separation of 11S and 2S proteins (Dalgalarrondo et al., 1986), and cation-exchange chromatography (CEC) to separate 2S isoforms (Monsalve and Rodriguez, 1990). More recently a new procedure for fractional enrichment and purification of plant (e.g. soy beans) and milk-based proteins was demonstrated by the use of volatile electrolytes for isoelectric precipitations (Hofland et al., 2003; Golubovic et al., 2005). Rapeseed protein meal has two major classes of seed storage proteins: 12S globulin (cruciferin), which represents 25–65% of its protein content (Raab et al., 1992) and 2S albumin (napin). In addition to these two proteins it also contains some minor proteins of interest, such as thionins, trypsin inhibitors and a lipid transfer protein (LTP). From the structure and physicochemical properties of these proteins it is evident that cruciferin shows very different characteristics from other proteins including: high molecular weight, neutral isoelectric point pI; but aggregation could occur during the purification process. On the contrary, napin and LTP show rather close characteristics in their molecular weights and pIs, which could complicate the purification process. Separation of napin and cruciferin can be obtained by a combination of nanofiltration in combination with several chromatographic steps of cationexchange, size-exclusion and hydrophobic interaction chromatography on process scales.
4.4 Recent developments in process chromatography Figure 4.8 shows some of the areas where there is potential for development in process scale chromatography. The innovations can be categorized in several fields, materials, modes of operation, and equipment. In the subsequent subsections some of the innovation in these areas will be highlighted. 4.4.1 Structured matrices and monoliths Monoliths for enhancing reactions have been around for several decades. For example, PolyHipe is a highly porous macrocellular cross-linked styrene–divinylbenzene (DVB) copolymer, prepared by polymerization of a high internal phase emulsion (HIPE) of water droplets dispersed in a styrene/DVB continuous phase (Ottens et al., 2000). This type of monolith © Woodhead Publishing Limited, 2010
Advances in process chromatography and applications 129 ∑ Materials – optimization of selectivity, MT, and hydrodynamics, magnetic beads, monoliths, structured matrices ∑ Mode of operation – exploiting non-linear effects of coupled multi-component equilibria ∑
Equipment – CPC: centrifugal partitioning chromatography – Radial flow: annular chromatography – SMB (counter-current chromatography) – EBA (expanded bed adsorption)
∑ Chromatographic reactors – coupled reaction – separation
I Feed A,B
Waste
II III IV
Solids B Raffinate A Extract
Desorbent
Fig. 4.8 Areas of innovation in chromatography and a schematic diagram of a simulated moving bed.
was used to enhance a multiphase reaction, but had structural properties that also made it suitable for high capacity, low-pressure drop chromatographic separations. Structured monoliths from Corning that have been used in chemical conversions in the work of Kreuzer et al. (2005) would also suit chromatographic separations purposes. Pictures of monoliths are given in Fig. 4.9. For more than a century, chromatography practitioners have been separating the components of chemical mixtures by using columns packed with various types of particulate matter. Monoliths are a relatively new class of stationary phases, completely different from conventional stationary phases. A good overview of this relatively new stationary chromatographic phase in the biotechnology area is given by Jungbauer and Hahn (2004). The material is cast into a chromatography column as a continuous block of one piece interlaced with channels (Jungbauer et al., 2002a). The ramified channels do not have dead ends. Owing to this structure, the transport of the solute to the surface is solely by convection instead of diffusion as observed in conventional media (Hahn and Jungbauer, 2000; Hahn et al., 2002;). These nontraditional column materials have recently been commercialized and are being recommended by manufacturers and users for the enhanced speed and thoroughness with which they can separate complex mixtures of biological molecules. The most important characteristics of monolithic media are the excellent mass transfer properties and the low pressure drops. The large channel diameter makes monoliths excellent stationary phase materials for chromatographic separations. The first monoliths were developed by Hjerten et al. (Hjerten and Liao, 1988; Hjerten et al., 1989, 1992) and Tennikova et al. (1990). Hjerten and coworkers compressed polyacrylamide gels and observed excellent resolution. Scale up was initially difficult but recently was successfully performed (Podgornik et al., 2000). Another important development is in the use of monoliths as a support for solid-phase synthesis (Jungbauer et al., 2002c). An interesting application is the direct synthesis of peptides on © Woodhead Publishing Limited, 2010
130 Separation, extraction and concentration processes
(a)
(b)
Fig. 4.9 Photographs of monoliths: (a) SEM photograph (15 kV; scale, 1 cm = 32.7 mm) of PolyHipe-type X20PV90 on an ISI-DS-130 SEM apparatus (Ottens et al., 2000); (b) Corning monoliths, as used in catalytic multiphase conversions by Kreutzer et al. (2005), with potential for use in large-scale food, beverage and nutraceutical purification.
monolithic columns. Because the synthesis is performed on polymethacrylate monoliths, the directly grown peptide can be used as an affinity ligand without any further treatment (Pflegerl et al., 2002a; 2002b; 2002c). This strategy provides an excellent screening platform for affinity ligands. The screening of the ligand can be performed by microtitration and the same resin can be used for screening and large-scale separation. Monoliths are also interesting supports for enzyme reactors. In packed-bed reactors, the efficiency is often limited by pore diffusion. The high porosity of monoliths ensures that the enzymatic process is not mass transfer limited (Jungbauer et al., 2002b). © Woodhead Publishing Limited, 2010
Advances in process chromatography and applications 131 A relatively new stationary phase is the combination of chromatographic resin beads and a membrane. Mixed matrix membranes (MMMs), which incorporate adsorptive particles during membrane casting, can be prepared simply and have performances that are competitive with other membrane chromatography materials. The application of MMM chromatography for fractionation of b-lactoglobulin from bovine whey was recently shown (Saufi and Fee, 2009). The dynamic binding capacity of b-lactoglobulin in whey solution was about 80 mg g–1 membrane (24 mg mL–1 of MMM), which is promising for whey fractionation using this technology. 4.4.2 Novel ligands Affinity separations are normally very expensive; therefore, application in the food industry requires either very high value products or a lower ligand cost. Affinity ligands are the basic moieties on which the various affinity chromatographic separations run. Affinity chromatography has been used widely in biomedical research and biotechnology (Wilchek and Chaiken, 2000). It is based on molecular recognition where one recognition partner is immobilized on a base matrix, and soluble target molecules can be retained from a crude mixture. The target molecule can then be released and recovered in a functional form. The basis of elution is to reduce the affinity between immobilized ligand and analyte. This is most often accomplished by use of a bond-breaking buffer (i.e., by changing pH, ionic strength, and solvent) or by use of competitive elution (Firer, 2001). In either instance, the diffusion of immobilized ligand is always strictly limited, and therefore, the local concentration of immobilized ligand is constant. The environment surrounding the immobilized ligand changes to modulate the ligand–analyte affinity. Affinity chromatography is potentially the most selective method for protein purification. The technique has the purification power to eliminate steps, increase yields and thereby improve process economics. Although affinity chromatography is used extensively on a laboratory scale, its widespread acceptance has been limited on the preparative scale because of the high cost of the affinity ligands and their biological and chemical instability (Lowe et al., 2001). Only recently, the development of new methods for screening, selection and design of stable synthetic ligands, has opened the opportunity of exploitation of such materials on a large-scale (Ladner and Ley, 2001). The rapid growth of bioinformatics and molecular docking techniques and the introduction of combinatorial methods for systematic generation and screening of large numbers of novel compounds, has made feasible the rapid and efficient generation of ligands for affinity chromatography (Clonis, 1990; Labrou, 2002; 2003; Labrou et al., 2004; Spalding, 1991). Immobilized lectin affinity chromatography (Endo, 1996; Kobata, 1994) using Ricinus communis agglutinin (RCA) has been used for separation of several glycopeptides and oligosaccharides (Merkle and Cummings, 1987;
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132 Separation, extraction and concentration processes Shinohara et al., 1987). RCA binds specifically to nonreducing-end galactose residues (Wei and Koh, 1978). Competitive elution using lactose, a hapten sugar, is generally used to elute the target molecule from the RCA immobilized column. Separation of bioactive peptides from defatted sunflower meal using immobilized lung extracts or immobilized angiotensin converting enzyme (ACE) (Megıas et al., 2004; 2006a; 2006b; 2007a; 2007b) has been used. Peptides that have angiotensin converting enzyme (ACE)-inhibitory activity are of great interest because of their potential antihypertensive effect, and have been found in hydrolysates of plant and animal origin (Murray and FitzGerald, 2007; Ottens et al., 2006a; 2006b). These are a few examples of using novel ligands in the separation of several valuable products from food sources. Enrichment by affinity chromatography generally allows further purification using classical chromatographic techniques. 4.4.3 Mode of operation Counter-current chromatography: simulated moving bed In conventional chromatography, the resin is fixed in a column, and the liquid flows through the column. As in extraction and distillation processes, the counter-current movement of the two phases improves the efficiency of the separation process, by establishing a higher driving force for mass transfer. A typical moving bed system is shown in Fig. 4.10b. It is equivalent to a distillation column. The liquid flows from section I to section IV, the sorbent moves from section IV to section I. A feed mixture (F) is introduced in the middle of the system. It is fractionated in an extract stream (E) containing the more retained component(s), and a raffinate stream (R), containing the less retained component(s). Similar to distillation, the movement of the components depends on the flow rates of the liquid and solid phases and the equilibrium distribution coefficient over both phases. Commonly, a four section SMB is used, the middle sections II and III serving to achieve the actual separation and section I cleaning the sorbent for reuse while section IV cleans the liquid for (partial) reuse. A real continuous movement of the solid phase is difficult to achieve in chromatography without losing resolution, owing to attrition of the solid phase. In practice, the movement of the solid phase is ‘simulated’ by periodic switching of a carousel of fixed chromatographic columns (Fig. 4.10a). The efficient use of the resin and liquid (desorbent D) is an important contribution to the efficiency of the SMB compared with that of fixed bed (FB) chromatography. Normally a reduction by a factor of 2–10 in resin and buffer use per amount of product produced can be achieved in an SMB compared with a FB. Chiral separation was the main driving force for the establishment of SMB separations in the fine chemicals and pharmaceutical industry over the past two decades (e.g. Cavoy et al., 1997; Pais et al., 1997; Rodriques and Pais, 2004). It is now a well-known and regularly applied chiral binary separation
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Advances in process chromatography and applications 133 Feed Liquid flow cycle
Liquid flow direction
Raffinate
Extract Liquid flow direction
Desorbent
(a) Sorbent recycle q4
q3
q2
IV c4
c3
II c2 c¢2
R
q0
q1
III
cF F
I c1
c0
D
E
N species: 11 N unknowns (c, q)0, …, 4 + c¢2 11 N independent equations (all linear)
4N concentration ratio 4N overall mass balances 3N balances: feed, desorbent, sorbent [A] x = b (b)
Fig. 4.10 (a) Schematic of the concept and operation of continuous multi-column simulated moving bed chromatography. Different tints indicate the extent of concentration of the two solutes. (b) A simple steady-state model for mathematical modeling of four-zone linear SMB chromatography: q solute solid-phase concentration, c solute liquid-phase concentration, R the raffinate stream, F the feed stream, E the extract stream, D the desorbent stream. The different zones are indicated by I–IV and 1–4.
technique. SMB is also a well-studied operation. However, some areas of application need to be developed. Separations of biologicals (Houwing, 2003; Houwing et al., 2002, 2003a; 2003b; Paredes et al., 2005) as well as multi-component (MC) separations in SMBs (e.g. Abel et al., 2004; Wang and Ching, 2005) are important research fields with a large industrial relevance. Multi-component (MC) separations can be performed in SMBs and some studies are devoted to it, mainly by serial SMB systems. Multiplezone (or section) SMBs have been reported (up to nine zones, Wooley © Woodhead Publishing Limited, 2010
134 Separation, extraction and concentration processes et al., 1998). Five-zone SMBs are being investigated and design criteria have been presented, for example, by Kim et al. (2003b). Wang and Ching (2005) presented a straightforward extension of a four-section SMB to a five-section SMB. An extra raffinate or extract port was added to the standard four-section SMB to perform a ternary separation. The pros and cons of two raffinate or two extract streams were discussed. However, in all these cases, a single component was purified from a three-component mixture. An equal or larger volume of industrial separations, particularly in the food industry, concern enrichment of a (MC) product in a particular functionality, for instance from a health or nutritional point of view. When using chromatography, the functionality can be associated to a particular MC fraction of an (also) MC feedstock. Developing and optimizing an efficient separation has to deal with the inherent MC nature of the feedstock and product(s). Also in chromatographic separations in food and biotechnology, consumption of costly resin and salty buffers is high, which are reasons to use SMB technology for these applications as well. As yet, few biotechnology applications of SMB technology have been described in open literature. The few examples found mainly concern small biomolecules, such as amino acids (e.g. Maki et al., 1987; Van Walsum and Thompson, 1997; Wu et al., 1998) disaccharides (Geisser et al., 2005) polypeptides (e.g. Mun et al., 2003) but also protein separations and purifications (e.g. Adachi, 1994; Gottschlich et al., 1996; Gottschlich and Kasche, 1997; Hashimoto et al., 1988; Horneman et al., 2006, 2007a; Houwing, 2003; Houwing et al., 2002, 2003a; Huang et al., 1988), plasmid purifications (Paredes et al., 2005) and viral clearance (Horneman et al., 2007b). In biotechnology, a whole range of chromatographic techniques is available for separation of proteins. SMB is applicable to any of these systems. Considering size exclusion; several forms have been investigated: separation of dilute mixtures (i.e., the linear isotherm case), surfactant-aided gel filtration (Horneman et al., 2004, 2006, 2007a; 2007b), separation of concentrated mixtures (the non-linear isotherm case) (Houwing, 2003), and separation of multicomponent mixtures (Ottens et al., 2006a). In the latter work, the technological and economic feasibility of SMB technology for multi-component product (and feedstock) separations is described, using the fractionation of peptides from protein hydrolysates as functional food ingredients as an example. The impact on the environment was also assessed. The model system consisted of an enzymatic hydrolysate of casein, which was fractionated using gel filtration (size-exclusion chromatography). The technical feasibility of isolating peptides, capable of inhibiting the ‘angiotensinconverting enzyme’ (ACE), from a casein hydrolysate by means of size-exclusion chromatography was investigated experimentally. In another study, the successful separation of the disaccharide lactose from a complex mixture of human milk oligosaccharides (HMOS) with the continuous chromatography of SMB technique is described (Geisser et al., 2005).
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Advances in process chromatography and applications 135 Expanded bed adsorption A major concern when operating a chromatographic column is fouling and, clogging of the column causing high pressure drop and preferential flow, or channeling, and thereby loss of capacity. To overcome this problem a technique was developed over recent decades that overcomes this problem, by operating the adsorption and washing step in the fluidized mode. Liquid enters the column from below at sufficient velocity to drag the sedimented solid stationary material with it, thus creating a fluidized or expanded bed. This mode of operation allows processing of crude feed streams with a high load of particulates, and can be used as a combined solid–liquid separation step and an initial capture step (e.g. Hubbuch et al., 2006). A recent application of expanded bed adsorption (EBA) in the food industry is the recovery fractionation and purification of valuable proteins from potato juice (Løkra et al., 2009). Horneman et al. (2008) used EBA in SMB mode to optimally recover and purify potato proteins to obtain an economically viable process.
4.5 Process control in chromatography In order to have an optimal separation at high productivity, yield and purity, it is necessary to operate under the appropriate conditions. The conditions may change owing to fouling or loss of capacity, a fluctuating feed composition or the inherent unsteady state of the operational procedure (i.e. with SMB chromatography). To operate at optimal conditions, necessitates control of the operation of the industrial-scale chromatography column(s). For SMB a safe regime is usually 10% below the optimal boundary. Proper control strategies may allow higher productivity, while maintaining the production constraints of purity and yield (Grossmann et al., 2008). With the availability of massive, cheap computer power and detailed mathematical models (as outlined in section 4.2.3), process scale chromatography will benefit more and more in the near future from proper control of the operation of the unit.
4.6 Future trends 4.6.1 High-value nutraceuticals Functional foods and nutraceuticals provide an opportunity to improve health while reducing health care costs. Functional foods involve the broad class of prebiotics and probiotics. Probiotics are viable microbial dietary supplements that influence the host and have a beneficial effect in the gastro-intestinal tract (Jardine, 2009; Salminen et al., 1998), whereas prebiotics are nondigestible food ingredients that benefit the host organism by stimulating the growth or activity of one or limited number of bacteria in the colon (Gibson and Roberfroid, 1995; Jardine, 2009). © Woodhead Publishing Limited, 2010
136 Separation, extraction and concentration processes Nutraceuticals refers to extracts of foods claimed to have a medicinal effect on human health. An important trend in food technology is the production of these ingredients from sources such as aromatic plants or spices. Among the nutraceuticals, antioxidants receive much attention in the food industry (Madhavi et al., 1996), not only as preservatives in food products to prevent or retard oxidation of fats and oils, but also because of their beneficial effects on human health. Supercritical-fluid chromatography (SFC) is one of the widely used process options in nutraceutical separations. Preparative-scale SFC is an environmentally clean technology whose main advantage, compared with preparative LC, is the easy recovery of the isolated compounds by a simple decompression of the supercritical fluid (Coleman et al., 1999). Fish oil preparation and separation of polyunsaturated fatty acid esters such as docosahexaenoic acid (DHA) and eicosapentaenoic acid (EPA) (Alkio et al., 2000) is a well-documented example of preparative SFC and has been used in nutraceutical production since the early 1990s. Novel methods such as continuous membrane chromatography reactor system (CMCRS) are used for the synthesis and separation of galacto oligosaccharides (GOS) (Engela et al., 2008). Membrane separations using ion-exchange chromatography is another widely used technique for functional components from milk sources (Goodall et al., 2008). 4.6.2 Good manufacturing practice (GMP), quality assurance (QA) and quality control (QC) and Food and Drug Administration (FDA) regulation Good manufacturing practice (GMP) is a term that is recognized worldwide for the control and management of manufacturing and quality control of foods, pharmaceutical products, and medical devices. GMPs are guidelines that outline the aspects of production that affect the quality of a product. The guiding principle of GMP is that quality is built into a product, and not just tested into a product. Therefore, the assurance is that the product not only meets the final specifications, but that it has been made by the same procedures, called standard operating procedures (SOP), under the same conditions each and every time it is made. In the food industry GMPs address all factors involved in a manufacturing process: personnel, building, premises, machinery and apparatus, documentation, and quality control, that generally affect product quality or impact quality monitoring. Besides these factors, an important theme that is part of GMP is validation. It is that part of GMP that ensures that facility systems, equipment, processes, and test procedures are under control and therefore consistently produce quality product. This is a special quality assurance method, which is explicitly required by GMP codes and mandated by law. GMPs are enforced in the United States by the FDA (http://www.fda.gov); within the European Union, GMP inspections are performed by National
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Advances in process chromatography and applications 137 Regulatory Agencies, e.g. GMP inspections are performed in the United Kingdom by the Medicines and Healthcare products Regulatory Agency (MHRA)(http://www.mhra.gov.uk), in the Republic of Korea (South Korea) by the Korea Food and Drug Administration (KFDA), and by similar national organizations worldwide. 4.6.3 Process analytical techniques (PAT) and control Process analytical techniques (PAT) are now being implemented in the pharmaceutical industry to monitor and control proper operation of these processes. Also in chromatographic fractionating and purification in the food industry such an application is seen. It allows for GMP, QC, better operation, better control, higher productivity and on spec production/higher number of on spec batches.
4.7 Conclusions Process chromatography is being applied in food, beverage and nutraceutical processing. Chromatography has a high resolving power which can give highpurity food products. With the advent of the nutraceutical field, also more expensive forms of chromatography, based on affinity, may find application, as the potential higher sales revenues balance the higher operating costs. Chromatography is a flexible purification method that is likely to play a role in food, beverage and nutraceutical processing for years to come.
4.8 Sources of further information and advice ∑
Ganetsos, G., Barker, P.E. Preparative and production scale chromatography. Marcel Dekker, Inc., 1992. ∑ Guiochon, G., Shirazi, S.G., Katti, A.M. Fundamentals of preparative and nonlinear chromatography. Academic Press, Inc., 2nd Edition, 1994. ∑ Sofer, G., Hagel, L. Handbook of process chromatography – a guide to optimization, scale-up and validation; Academic Press, London, 3rd Edition, 2001.
4.9 List of abbreviations ACE CEC CMCRS
angiotensin converting enzyme cation-exchange chromatography continuous-membrane chromatography reactor © Woodhead Publishing Limited, 2010
138 Separation, extraction and concentration processes CSS DHA DVB EBA EPA FB FDA GMP GOS HIPE HMOS IEC IgG LEC LF LP LTP MC MHRA MMM MT PAT PDE QA QC RCA RP-HPLC SEC SFC SMB SOP
combined stabilization system docosahexaenoic acid divinylbenzene expanded bed adsorption eicosapentaenoic acid fixed bed Food and Drug Administration good manufacturing practice galacto oligosaccharides high internal phase emulsion human milk oligosaccharides ion-exchange chromatography immunoglobulin G ligand-exchange chromatography lactoferrin lactoperoxidase lipid transfer protein multi-component Medicines and Healthcare products Regulatory Agency mixed matrix membranes mass transfer process analytical techniques partial differential equation quality assurance quality control ricinus communis agglutin reversed-phase high-performance liquid chromatography size-exclusion chromatography supercritical-fluid chromatography simulated moving bed standard operating procedure
4.10 References Abel, S., Babler, M.U., Arpagaus, C., Mazzotti, M., Stadler, J., (2004) Two-fraction and three-fraction continuous simulated moving bed separation of nucleosides, J Chromatogr A, 1043(2): 201. Adachi, S., (1994) Simulated moving bed chromatography for continuous separation of two components and its application to bioreactors, J Chromatogr, 658: 271. Alkio, M., Gonzales, C., Jäntti, M., Aaltonen, O., (2000) Purification of polyunsaturated fatty acid esters from tuna oil with supercritical fluid chromatography, J Am Oil Chem Soc, 77: 315–321. Al-Mashikhi, S.A., Li-Chan, E., Nakai, S., (1988) Separation of immunoglobulins and lactoferrin from cheese whey by chelating chromatography, J Dairy Sci, 71: 1747–1755.
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Advances in process chromatography and applications 145 Noh, K.M., Imm, J.Y., (2005) One-step separation of lysozyme by reverse micelles formed by the cationic surfactant, cetyldimethylammonium bromide, Food Chem., 93: 95–101. Noppe, W., Haezebrouck, P., Hanssens, I., De Ouyper, M., (1999) A simplified purification procedure of alpha-lactalbumin from milk using Ca 2+-dependent adsorption in hydrophobic expanded bed chromatography, Bioseparation, 8(1/5), 153–158. Ottens, M., Houwing, J., van Hateren, S.H., van Baalen, T. van der Wielen, L.A.M., (2006a) Multi-component fractionation in SMB chromatography for the purification of active fractions from protein hydrolysates, Food Bioprod. Process., 84(C1), 59. Ottens, M., Leene, G., Beenackers, A.A.C.M., Cameron, N., Sherrington, D.C., (2000) PolyHipe: a new polymeric support for heterogeneous catalytic reactions: kinetics of hydration of cyclohexene in two- and three-phase systems over a strongly acidic sulfonated PolyHipe, Ind. Eng. Chem. Res., 39: 259–266. Ottens, M., Wesselingh, J.A., van der Wielen, L.A.M., (2006b) Downstream processing in biotechnology, Chapter 9 in Basic biotechnology, edited by Ratledge and Kristiansen, 3rd edition, Cambridge University Press. Owen, R.O., Chase, H.A., (1997) Direct purification of lysozyme using continuous counter-current expanded bed adsorption, J. Chromatogr. A, 757: 41–49. Paananen, H., (1996) Trends in the chromatographic separation of molasses, SPRI 1996 Workshop on Separation Processes in the Sugar Industry, New Orleans, USA. Paananen, H., Kuisma, J., (1999) Multiple value-added products using the FAST separation technology, International Conference on Value-Added Products for the Sugar Industry, Baton Rouge, USA. Paillat, M.C.P., (1999) Different industrial applications of continuous chromatography in the sugar industry and for the production of sugar derivatives, Detmold Starch Convention, Detmold, Germany. Pais, L.S., Loureiro, J.M., Rodrigues, A.E., (1997) Separation of 1,10-bi-2-naphthol enantiomers by continuous chromatography in simulated moving bed, CES, 52(2): 245. Paredes, G., Makart, S., Stadler, J., Mazzotti, M., (2005) Simulated moving bed operation for size exclusion plasmid purification, CET, 28(11): 1335. Pearce, R.S., Houlston, C.E., Atherton, K.M., Rixon, J.E., Harrison, P., Hugues, M.A., Dunn, M.A., (1998) Plant Physiol., 117: 787. Pflegerl, K., Podgornik, A., Berger, E., Jungbauer, A., (2002b) Screening for peptide ligands on CIM monoliths, Biotechnol. Bioeng., 79: 733–740. Pflegerl, K., Podgornik, A., Schallaun, E., Jungbauer, A., (2002c) Direct synthesis of peptides on CIM monolithic columns for affinity chromatography, J. Combinat. Chem., 4: 33–37. Pflegerl, K., Hahn, R., Berger, E., Jungbauer, A., (2002a) Mutational analysis of a blood coagulation factor VIII-binding peptide, J. Peptide Res., 59: 174–182. Plate, K., Beutel, S., Buchholz, H., Demmer, W., Fischerfruhholz, S., Reif, O., Ulber, R., Scheper, T. (2006) Isolation of lactoferrin, lactoperoxidase and enzymatically prepared lactoferrin from proteolytic digestion of bovine lactoferrin using membrane adsorptive chromatography, J. Chromatogr. A, 1117, 81–86. Podgornik, A, Barut, M., Strancar, A., Josic, D., Koloini, T., (2000) Construction of large-volume monolithic columns, Anal. Chem., 15: 5693. Raab, B., Leman, H., Schwenke, K.D., Kozlowska, H., (1992) Comparative study of the protein patterns of some rapeseed (Brassica napus L.) varieties by means of polyacrylamide gel electrophoresis and high-performance liquid chromatography, Nahrung, 36: 239–247. Raab, B., Schwenke, K.D., (1984) Simplified isolation procedure for the 12 S globulin and the albumin fraction from rapeseed (Brassica napus L.), Nahrung, 8: 863–866. Rehmanji, M., Gopal, C., Mola, A., (2000) Tech. Q. Master Brew. Assoc. Am., 39: 24–28, A novel stabilization of beer with Polyclar Brewbrite.
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146 Separation, extraction and concentration processes Rito-Palomares, M., Miguel H., (1998) Influence of system and process parameters on partitioning of cheese whey proteins in aqueous two-phase systems, J. Chromatogr. B, 711, 81. Rodrigues, A.E., Pais, L.S., (2004) Design of SMB chiral separations using the concept of separation volume, Sep. Purif. Technol., 39(2): 245. Rousset. F., (1997) New developments in chromatographic separation of beet molasses. British Sugar EuroTechLink 97. Roy, I., Rao, M.V.S., Gupta, M.N., (2003) Purification of lysozyme from other hen’segg-white proteins using metal-affinity precipitation, Biotechnol. Appl. Biochem., 37: 9–14. Ruckenstein, E., Zeng, X.F., (1997) Macroporous chitin affinity membranes for lysozyme separation, Biotechnol. Bioeng., 56: 610–617. Safarik, I., Safarikova, M., (2003) Batch isolation of hen egg white lysozyme with magnetic chitin, J. Biochem. Biophys. Methods, 27: 327–330. Safarik, I., Safarikova, M., (2004) Magnetic techniques for the isolation and purification of proteins and peptides, Biomagn. Res. Technol., 2: 7. Salminen, S., Bouley, C., Boutron-Ruaultm M.C., Cummings, J.H., Franck, A., Gibson, G.R., Isolauri, E., Moreau, M.G., Roberfroid, M., Rowland, I.R. (1998) Functional food science and gastrointestinal physiology and function. Br. J. Nutr., 80(suppl): S147–71. Saufi, S.M., Fee, C.J., (2009) Fractionation of b-lactoglobulin from whey by mixed matrix membrane ion exchange chromatography, Biotechnol. Bioeng., 103: 138–147. Shinohara, Y., Kim, F., Shimizu, M., Goto, M., Tosu, M., Hasegawa, Y., (1994) Kinetic measurement of the interaction between an oligosaccharide and lectins by a biosensor based on surface plasmon resonance, Eur. J. Biochem., 223: 189–194. Spalding, B.J., ((1991) Downstream processing: key to slashing production costs 100 fold, Bio/Technology, 9: 229. Stahl, B., Thurl, S., Zeng, J., Karas, M., Hillenkamp, F., Steup, M., Sawatzki, G., (1994) Oligosaccharides from human milk as revealed by matrix-assisted laser desorption/ ionization mass, Anal. Biochem., 223: 218. Steffenson, M., Westerlund, D., (1996) J. Chromatogr. A, 720: 127–136. Su, C.K., Chiang, B.H., (2006) Partitioning and purification of lysozyme from chicken egg white using aqueous two-phase system, Process Biochem., 41: 257–263. Tejeda-Mansir, A., Montesinos, R.M., Magana-Plaza, I., Guzman, R., (2003) Breakthrough performance of stacks of dye–cellulosic fabric in affinity chromatography of lysozyme, Bioprocess Biosyst. Eng., 25: 235–242. Tennikova, T., Svec, F., Belenkii, B.G., (1990) High-performance membrane chromatography. A novel method of protein separation, J. Liq. Chromatogr. 13: 63–70. Thanh, V.H., Shibasaki, K., (1976) Proteins of soybean seeds. A straightforward fractionation and their characterisation, J. Agric. Food Chem., 24: 1117–1121. Thurl, S., Henker, J., Taut, H., Tovar, K., Sawatzki, G., (1993) Variations of neutral oligosaccharides and lactose in human milk during the feeding, Z. Ernahrungswiss., 32: 262–269. Thurl, S., Offermanns, J., Müller-Werner, B., Sawatzki, G., (1991) Determination of neutral oligosaccharide fractions from human milk by gel permeation chromatography, J. Chromatogr., 568: 291. Tomita, M., Wakabayashi, H., Yamauchi, K., Teraguchi, S., Hayasawa, H., (2002) Bovine lactoferrin and lactoferricin derived from milk: production and applications, Biochem. Cell Biol., 80(1): 109–112. Torre, M., Cohen, M.E., (1996) Perfusion liquid chromatography of whey proteins, J. Chromatogr. A, 729: 99–111. Van Walsum, H.J., Thompson, M.C., (1997) Simulated moving bed in the production of lysine, J. Biotechnol., 59: 127. Visser, S., Slangen, C.J., Rollema, H.S., (1991) Phenotyping of bovine milk proteins by
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Advances in process chromatography and applications 147 reversed-phase high-performance liquid chromatography, J. Chromatogr., 548(1–2): 361–370. Wang, X., Ching, C.B., (2005) Chiral separation of b-blocker drug (nadolol) by five-zone simulated moving bed chromatography, CES, 60: 1337–1347. Wei, C.H., Koh, C.J., (1978) Crystallographic characterization of a principal non-toxic lectin from seeds of Ricinus communis, Mol. Biol., 123, 707–711. Wilchek, M., Chaiken, I., (2000) An overview of affinity chromatography, Methods Mol. Biol. 147, 1–6. Wooley, R., Ma, Z., Wang, N.-H.L., (1998) A nine-zone simulating moving bed for the recovery of glucose and xylose from biomass hydrolyzate, IECR, 37: 3699–3709. Wu, D.J., Xie, Y., Wang, N.H.L., (1998) Design of simulated moving bed chromatography for amino acid separations, IECR, 37: 4023. Xu, Y., Sleigh, R., Hourigan, J., Johnson, R., (2000) Separation of bovine immunoglobulin G and glycomacropeptide from dairy whey, Process Biochem., 36: 393–399. Yamada, H., Fukumura, T., Ito, Y., Imoto, T., (1985) Chitin-coated celite as an affinity adsorbent for high performance liquid chromatography of lysozyme, Anal. Biochem., 146: 71–74. Yamasaki, N., Eto, T., (1981) A novel adsorbent for affinity chromatography of lysozyme, Agric. Biol. Chem., 45: 2939–2941. Ye, X., Yoshida, S., Ng, T.B., (2000) Isolation of lactoperoxidase, lactoferrin, a-lactalbumin, b-lactoglobulin B and b-lactoglobulin A from bovine rennet whey using ion exchange chromatography, Int. J. Biochem. Cell Biol., 32: 1143. Yoshida, S., (1988) Isolation of some minor milk proteins, distributed in acid whey from approximately 100,000 to 250,000 daltons of particle size, J. Dairy Sci., 71(1): 1–9. Yoshikawa, K., Okamura, M., Inokuchi, M., Sakuragawa, A., (2007) Ion chromatographic determination of organic acids in food samples using a permanent coating graphite carbon column, Talanta, 72: 305–309. Zydney, A.L., (1998) Protein separations using membrane filtration: new opportunities for whey fractionation, Int. Dairy J., 8, 243–250.
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5 Novel adsorbents and approaches for nutraceutical separation B. W. Woonton, CSIRO Food and Nutritional Sciences, Australia and G. W. Smithers, Food Industry Consultant, Australia
Abstract: In this chapter, an overview is provided of novel chromatographic adsorbents that are perhaps on or just over the horizon but hold promise as the basis for future industrial technologies in the cost-effective extraction of nutraceuticals from agri-food streams. Such adsorbents include molecular imprinted polymers, organic monoliths, stimuli-responsive resins, mesoporous molecular sieves, peptide affinity ligands, and membrane adsorbers. These adsorbents have the potential to improve specificity, selectivity, simplicity, robustness, and productivity, and to reduce the environmental impact of nutraceutical manufacture. Key words: nutraceuticals, separation, molecular imprinted polymers, monoliths, stimuli-responsive resins, molecular sieves, peptide affinity ligands, membrane adsorbers.
5.1 Introduction Nutraceuticals have entered the mainstream of foods and beverages. By definition, these components derived from agricultural and food raw materials have a positive influence on human health through a specific prophylactic or therapeutic physiological effect against chronic disease when consumed (Hardy, 2000). Nutraceuticals have usually been delivered to the consumer in a medicinal vehicle such as a tablet, capsule or in powdered form. More recently, the functional foods market, driven by consumer demands for healthpromoting foods and drinks, has seen nutraceuticals enter the food industry by way of the active component(s) in ‘functional’ varieties of everyday food and beverage products like yoghurt, bread, margarine, juice, and even water. This market is becoming a lucrative outlet for new and novel nutraceuticals, and © Woodhead Publishing Limited, 2010
Novel adsorbents and approaches for nutraceutical separation 149 food companies are keen to capture a slice of this market. These companies are thus exploring new sources of nutraceuticals, and the means to isolate, characterize and substantiate the bioactivity of such components to support the development of successful functional foods and beverages. In an earlier chapter (Chapter 4), more traditional approaches (e.g. fixed and moving bed chromatography) to the extraction, separation and concentration of nutraceuticals have been addressed. Chromatography is a powerful technique for the cost-effective separation and isolation of nutraceuticals from agrifood streams, including food processing co-product streams. Developments in continuous chromatography have enabled the commercial separation and isolation of biologically active proteins from milk and whey. For example, continuous separation (CSEP) technology is being used to isolate valuable dairy-derived molecules, including nutraceuticals (De Silva et al., 2003). Central to the chromatographic process is the relevant adsorbent that is responsible for effecting the separation. In this chapter, we introduce novel chromatographic adsorbents and approaches for the isolation of nutraceuticals, those that are perhaps on/over the horizon but hold promise as future industrial technologies for the cost-effective isolation of nutraceuticals, and the potential to deliver on consumer demands for effective functional foods and beverages. This overview is not intended to be exhaustive or inclusive of all novel adsorbents and technologies, but rather to provide a summary of the most promising adsorbents and approaches, that with further research and development can be employed for large-scale nutraceutical manufacture. The chapter provides an account of developments in: (i) (ii) (iii) (iv) (v) (vi)
molecular imprinted polymers; organic monoliths; stimuli-responsive resins; mesoporous molecular sieves; peptide affinity ligands; and membrane adsorbers.
5.2 Molecular imprinted polymers and applications in the nutraceutical industry Molecular recognition is an important process in many biological systems, for example the specific recognition of an antigen by an antibody. Such recognition mechanisms have been utilized in the development of biological affinity-based separation media (e.g. mono- or poly-clonal antibodies and affinity peptide resins) (Tharakan et al., 1992). Molecular imprinted polymers (MIPs) differ from these biological affinity systems in that they are synthetic organic materials designed with cavities specific to a template molecule employed during synthesis. In effect, they are tailored synthetic affinity materials that can be used in a range of applications including the separation © Woodhead Publishing Limited, 2010
150 Separation, extraction and concentration processes and analysis of molecules, the sensing of molecules, as artificial enzymes, and as antibody mimetics. MIPs are worthy of discussion as they have the potential to reduce the complexity of nutraceutical separation processes, thus reducing processing costs and the environmental impact associated with such processes. The concept of molecular imprinting was first reported by Kikuchi et al. (1972) and Wulff & Sarhan (1972), and since these early reports, there has been a significant amount of research undertaken to improve MIPs using alternative polymer chemistries and synthesis methodologies. There are a large number of publications reporting the development and application of MIPs for a range of target molecules including small organic molecules such as pharmaceuticals, pesticides, amino acids, peptides, high molecular weight proteins, polypeptides, steroids, sugars and cells (Andersson et al., 1995; Bolisay et al., 2006; Bossi et al., 2007; Nicholls et al., 1995; Rachkov & Minoura, 2001; Sellergren, 1994; Shi et al., 1999; Steinke et al., 1995). 5.2.1 Synthesis The synthesis of MIPs involves the organization of a cross-linked polymer matrix around a template molecule, followed by removal of the template, leaving recognition sites complementary to the template or similar molecules (Fig. 5.1). There are three main steps involved in the synthesis of a MIP, and these include: (i) pre-arrangement of selected monomers (such as vinyl or acrylic) around the target molecule; (ii) polymerization of the monomers (e.g. radical polymerization) with the addition of a cross-linker; and (iii) extraction of the target molecule leaving binding sites that recognize the target molecule (Gupta & Kumar, 2008). Two different approaches may be employed to manufacture MIPs, these are covalent and non-covalent imprinting. Non-covalent imprinting is the most common and involves the interaction of the template molecule with the monomers via hydrogen bonding, ionic bonding, hydrophobic interactions and van der Waals forces (Mosbach, 1994). After polymerization and removal of the template, the functional groups of the polymer are positioned to rebind the target molecule using the same non-covalent interactions. The non-covalent approach can only be employed where template molecules have functional groups capable of strong non-covalent interactions with the monomers. However, this can be overcome by using the covalent imprinting method, which involves covalently attaching the template molecule to a monomer followed by polymerization, cleavage of the covalent bond to liberate the template, and removal of the template. The reactive groups are positioned within the polymer cavities to allow covalent bonds to be reformed with the target molecule and provide selectivity. As there is a relatively small number
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(i) Complexation Template
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Fig. 5.1 Schematic illustration of the method for preparing molecular imprinted polymers (Takeuchi & Haginaka, 1999). Reproduced with permission. Copyright Elsevier (1999).
Novel adsorbents and approaches for nutraceutical separation 151
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Functional monomer
152 Separation, extraction and concentration processes of suitable reversible covalent bonding reactions that can be employed, the covalent imprinting method has limited applications (Gupta & Kumar, 2008). MIPs are often manufactured using bulk polymerization methods to produce monolithic materials which can be used for in situ separations, or ground and sieved to produce particles suitable for column packing and separation applications (Haginaka, 2009). MIPs can also be manufactured using suspension, seed and dispersion/precipitation polymerization methods to produce micro- or nano-spheres (Haginaka, 2009; Kempe & Kempe, 2006; Ye & Mosbach, 2001). Using such suspension polymerization methods, MIPs can be generated in aqueous environments; however, hydrogen bonding between the template molecule and the functional monomer, and therefore MIP selectivity, can be reduced when using aqueous solvents (Haginaka, 2009). Suspension polymerization techniques in liquid perfluorocarbon or mineral oil do not reduce hydrogen bonding and electrostatic interactions and the resultant MIP microspheres have been shown to be selective for biologically active molecules (Perez-Moral & Mayes, 2006). MIPs are usually packed into columns of various dimensions to allow (a) sample preparation before analysis; (b) removal of valuable molecules; or (c) removal of undesirable molecules from various raw material streams. 5.2.2 Advantages and disadvantages The benefit of MIPs is their stability in various chemical and physical environments for an extended period of time without any loss in template recognition (Ciardelli et al., 2006; Kriz & Mosbach, 1995). In addition, they are physically robust, resistant to elevated temperatures and pressures, and inert towards acids, bases, metals ions and organic solvents. The reported limitations of MIPs include binding site heterogeneity and slow mass transfer, which subsequently leads to peak broadening and tailing. In addition, MIPs prepared by bulk polymerization often have irregular sized and shaped particles which can lead to poor chromatographic performance and large pressure drops (Turiel et al., 2005). These limitations may be overcome by: (i) optimizing the MIP loading solvent; (ii) optimizing the MIP washing regimen to remove contaminant molecules; and (iii) finer control over the polymerization reaction leading to more regularly shaped particles. 5.2.3 Applications The characteristics of MIPs make them ideal for use in separation applications including chromatography, capillary electrophoresis, solid phase extraction and membrane separations. They also have potential to be used as microreactors, in immunoassays, as antibody mimics and artificial enzymes. Owing to their © Woodhead Publishing Limited, 2010
Novel adsorbents and approaches for nutraceutical separation 153 selectivity and robustness, MIPs have been applied in the selective isolation, concentration and analysis of food-derived molecules such as theophylline, caffeine, catechin, cholesterol, glycyrrhetinic, theanine, a-tocopherol and flavonoids (Haginaka, 2009). For example, Puoci et al. (2007) developed a MIP and optimized conditions for the selective extraction and analysis of a-tocopherol from bay leaf (Fig. 5.2). This method recovered approximately 60% of the a-tocopherol with high purity as measured by HPLC. Many other MIPs have been reported for the separation and analysis of agricultural herbicides, pesticides and fungicides, toxins, antibiotics and other chemicals in food (Haginaka, 2009). The ability to produce MIPs for the selective isolation of proteins and peptides is improving (Bossi et al., 2007; Rachkov & Minoura, 2001; Shi et al., 1999), and being employed to selectively target bacteria through their surface expressed proteins (Xue et al., 2009). 5.2.4 Commercialization MIPs can be employed in analytical applications as solid-phase extraction media to selectively increase the concentration of analyte, decrease sample preparation time and increase method sensitivity. MIPs for the solid phase extraction and analysis of food contaminants such as antibacterials (chloramphenicol, nitroimidazoles, fluoroquinolones) and herbicides (triazines) have been manufactured by MIP Technologies AB (Lund, Sweden) and distributed by Sigma-Alridch (St Louis, USA). MIPs for solid phase extraction and analysis of vitamins in food such as riboflavin have recently become available (MIP Technologies AB, Sweden). There are few reports in the public domain verifying large-scale applications of MIPs in the food, pharmaceutical and chemical industries. However, given the reported benefits of MIPs including stability and specificity, with further research and development into large-scale manufacture of these adsorbents together with scaled applications, it is likely that they will find targeted use in the isolation of nutraceuticals.
5.3 Organic monoliths and applications in the nutraceutical industry In contrast to columns packed with small particles to facilitate separation (e.g. silica or agarose beads), monoliths are continuous solid phases which can be synthesized in situ in confined spaces such as columns, capillaries, and microfluidic and microchip devices. Monoliths are unique as they have an interconnected skeletal structure with flow-through paths that provide very high permeability to a mobile phase, high flow rates, and low back pressures during chromatography. These characteristics make monoliths promising for the large-scale cost-effective purification of nutraceuticals
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154 Separation, extraction and concentration processes 0.9 0.8
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Fig. 5.2 HPLC chromatograms of (a) crude bay leaf extract and (b) bay leaf extract after isolation and concentration of the a-tocopherol (a-TP) (Puoci et al., 2007). Reproduced with permission. Copyright Elsevier (2007). © Woodhead Publishing Limited, 2010
Novel adsorbents and approaches for nutraceutical separation 155 in the food industry. Monoliths can be synthesized using either inorganic materials such as silica or organic materials such as methacrylate polymers. As silica has poor stability and operational robustness during the caustic cleaning procedures employed in the food industry, this chapter focuses on the development and application of organic polymeric monoliths. 5.3.1 Synthesis Polymeric monoliths were first introduced in the late 1980s (Hjerten et al., 1989; Svec & Fréchet, 1992). Since then there have been several publications detailing developments in their manufacture and use in liquid chromatography, capillary electrochromatography and microscale separations (Nordborg & Hilder, 2009). Polymeric monoliths are synthesized by mixing an initiator, organic monomers (including a cross-linking monomer) such as styrene, methacrylate, acrylamide-based materials and pore forming solvents (porogens). The reagents are placed within a suitable mould, such as a column, capillary or channel of a microfluidic device and polymerized using radical polymerization initiated by either thermal treatment, ultraviolet, gamma or microwave radiation (Nordborg & Hilder, 2009). As polymerization proceeds, the polymer chains formed become insoluble in the polymerization mixture, phase separation takes place and the polymer chains precipitate out of solution, forming the monolith. The properties of the monolith synthesized, including the size and distribution of macro-, meso- and micropores are determined by the composition of the polymer mixture and the polymerization conditions. For example, the mode of initiation, type and amount of cross-linker, porogen, and the polymerization temperature. The large number of monomers that can be chosen to impart charged groups, other monomers, cross-linking agents and porogens results in a very large number of variations that can be examined and optimized to produce the desired monolith (Nordborg & Hilder, 2009). Organic monoliths can also be synthesized using sub-zero temperatures, where ice-crystals act as the pore-forming elements. After formation of the monolith, the ice crystals are allowed to thaw, the resulting water removed, leaving behind macropores between 10 and 100 mm in diameter. Such monoliths are known as cryogels (Arvidsson et al., 2002, 2003). Solid granules such as sodium sulfate can also be employed during the synthesis to improve bi-model pore formation and the hydrolytic characteristics of the resulting polymeric monoliths (Fig. 5.3) (Du et al., 2007). 5.3.2 Functionalization To improve the applicability of polymeric monoliths for separation applications, functionalization either during or after polymerization is required. In the first approach, the choice of monomers with charged functional groups results in material having charged groups throughout the monolith, including its surface.
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156 Separation, extraction and concentration processes (a)
Acc. V Spot Magn Det WD 20.0 3.0 2000¥ SE 9.3
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Fig. 5.3 SEM image of a glycidyl methacrylate-co-ethylene glycol dimethacrylate monolith prepared in the presence of the organic porogens dodecanol and cyclohexanol, (a) without and (b) with sodium sulfate granules (Du et al., 2007). Reproduced with permission. Copyright Elsevier (2007).
Examples of charged monomers that can be employed include methacrylic acid and ethylene glycol dimethacrylate. Optimization of monolith chemistry must take into account that the charged monomers participate in the polymerization process and will not all be available for separation interactions on the surface © Woodhead Publishing Limited, 2010
Novel adsorbents and approaches for nutraceutical separation 157 of the monolith. The approach of post-polymerization modification has been examined and reported to be advantageous as the number of variables is reduced and the monolith chemistry can be optimized followed by modification of the surface with the desired functionality. Monoliths can be modified to achieve separation based on ion-exchange, affinity recognition, reversedphase, and hydrophobic interaction. 5.3.3 Advantages and disadvantages Polymer monoliths have similar benefits to their polymeric particles. Their pH and pressure stability are two advantages that allow them to be employed in a range of environments where high flow rates may lead to large pressure drops across the bed. However, as monoliths are made from one piece of material with large interconnected pores (Fig. 5.3), diffusion mass transfer is minimized and convective mass transfer is maximized. As convective mass transfer is faster than diffusive transport, higher mobile phase flow rates can be employed and faster separation of molecules can be achieved. Breakthrough curves of lysozyme and immunoglobulin G using a polymethacrylate monolithic column (Connective Interaction Media, BIA Separations; Ljubljana, Slovenia) and a column packed with Source 30S particles (GE Healthcare Life Sciences Uppsala, Sweden) have been studied (Hahn et al., 2002). As expected, mass transfer in the particle packed column was very dependent on the linear velocity of the mobile phase, molecule size, and feed concentration, whereas mass transfer in the monolith was not significantly affected by velocity, molecule size or feed concentration (Hahn et al., 2002). A disadvantage of monoliths is their low surface area as a result of their high porosity. Thus, monoliths have been reported to be more suitable for the isolation of large biomolecules (e.g. proteins) rather than small molecules. Monoliths also have a tendency to be hydrophobic and thus adsorb proteins via both ion-exchange and hydrophobic interaction. This characteristic makes proteins difficult to elute from the monolith without solvent which can lead to protein denaturation (Nordborg & Hilder, 2009). Research to develop polymeric monoliths with less hydrophobic characteristics or modification of monolith surfaces with hydrophilic coatings is being undertaken to improve their effectiveness in the isolation of biological molecules. 5.3.4 Applications Monoliths have been employed to separate a range of food-derived biomolecules and nutraceuticals. These include enzymes such as pectin methylesterase, endo-polygalacturonase (Vovk et al., 2005; Vovk & Simonovska, 2007), lysozyme (Svec & Fréchet, 1995), and proteins such as trypsin inhibitor (Svec & Fréchet, 1995), bovine serum albumin (Du et al., 2007), lactoferrin (Adam et al., 2008), casein peptides (Gu et al., 2006), and virus particles
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158 Separation, extraction and concentration processes (Kramberger et al., 2004). Hilder et al. (2004) has also demonstrated that monoliths coated with quaternary amine-functionalized latex particles can effectively separate sugars (Fig. 5.4), and separate glucose and maltose from amylase-hydrolyzed corn starch. There are a number of studies indicating that monolithic adsorbents have potential for the preparative-scale separation and purification of biomolecules, such as immunoglobulins (Brne et al., 2007), a-lactalbumin, enzymes, plasmid DNA, viruses and cells (Jungbauer & Hahn, 2008). In addition, monolithic columns have been evaluated for their potential to isolate immunoglobulins using simulated moving bed chromatography (Pennings et al., 2008). 5.3.5 Commercialization Currently, a number of companies are manufacturing and distributing polymeric monolithic columns for analytical applications. These include BIA Separations (Villach, Austria), Dionex (Sunnyvale, USA), and BioRad (Hercules, USA). BIA Separations are one of only a few companies designing and supplying organic monoliths for both small- and large-scale purification. Their disk format monoliths are manufactured for small-scale separations using axial flow design characteristics, and their tube format monoliths are prepared for 3 1
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Fig. 5.4 Separation of carbohydrates within 10 min using an optimized anionexchange latex-coated polymeric monolithic capillary column. Peaks: (1) d(+) galactose, (2) d(+)glucose, (3) d(+)xylose, (4) d(+)mannose, (5) maltose, (6) d(–) fructose, (7) sucrose (Hilder et al., 2004). Reproduced with permission. Copyright Elsevier (2004).
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Novel adsorbents and approaches for nutraceutical separation 159 larger scale separation using radial flow design characteristics. The latter design is employed for large columns (e.g. up to 8 L) and is reported to allow for a significant increase in the linear velocity through the monolithic bed, without adversely affecting performance or pressure characteristics. Optimization of surface chemistry and adsorption capacity is a prerequisite for the application of organic monoliths in large-scale isolation of nutraceuticals, together with a determination of the economic benefit of using such adsorbents over traditional particle-packed columns. As manufacturing processes are optimized, and costs reduced, larger monolithic columns will deliver cost-efficient nutraceutical separation and isolation processes in the food industry.
5.4 Stimuli-responsive resins and applications in the nutraceutical industry Polymers that undergo a significant change in their structure and behavior (e.g. solubility) in response to an external physical, chemical or electrical stimulus are termed ‘stimuli-responsive’, ‘smart’ or ‘intelligent’ polymers. A number of polymers have this property and various stimuli, including temperature, pH and light can be applied to induce a change in polymer conformation and subsequent properties. Stimuli-responsive polymers with temperature responsiveness have been extensively studied and characterized. These polymers have been grafted onto chromatographic supports to produce stimuli-responsive resins. These novel resins have the potential to reduce solvent usage, processing costs and the environmental impact of nutraceutical isolation in the food industry. The best characterized temperature-responsive polymer is poly(Nisopropylacrylamide (poly-NIPAAm). This polymer differs from traditional polymers as its solubility decreases as the temperature of the poly-NIPAAm solution increases. When a critical temperature is reached (32 °C for polyNIPAAm), hydrogen bonding of water to the amide group is reduced and the poly-NIPAAm becomes unstable, contracts and enters a globular state (Fig. 5.5). The temperature at which this occurs is called the lower critical solution temperature (LCST) (Ayano & Kanazawa, 2006). The LCST of temperature-responsive polymers including poly-NIPAAm can be modified by incorporating hydrophilic or hydrophobic monomers into the polymer (Ayano & Kanazawa, 2006). Hydrophilic monomers generally increase the LCST of the polymer whereas incorporation of hydrophobic monomers generally decrease the LCST of the polymer. 5.4.1 Synthesis Temperature-responsive resins are synthesized by grafting temperatureresponsive polymers to the surface of solid chromatographic supports. Silica © Woodhead Publishing Limited, 2010
160 Separation, extraction and concentration processes
H 2O PNIPAAm
+ Stimulus – Temperature
Below LCST
Above LCST
Fig. 5.5 Coil to globule transition and subsequent solution turbidity change when poly-NIPAAm (PNIPAAm) is heated above or cooled below the lower critical solution temperature (LCST) (Ayano & Kanazawa, 2006). Reproduced with permission. Copyright Wiley-VCH Verlag GmbH & Co. KGaA.
has been the predominant support employed using radical co-polymerization of initiator immobilized silica or activated ester amine coupling to aminopropylmodified silica (Maharjan et al., 2008). Temperature-responsive co-polymers that have been grafted onto silica beads for chromatography include polyNIPAAm, poly-NIPAAm-co-butyl methacrylate, poly-NIPAAm-co-acrylic acid, and poly-NIPAAm-co-acrylic acid-co-N-tert-butylacrylamide (Maharjan et al., 2008). The use of silica matrices for the separation of nutraceuticals in the food industry is limited owing to their lack of operational robustness, including their instability at the high pH employed during cleaning processes (Maharjan et al., 2008). A temperature-responsive resin employing a porous crosslinked agarose support was recently reported. This resin was synthesized by functionalizing agarose beads with the initiator 4,4-azobis(4-cyanovaleric acid), followed by polymerization of N-isopropylacrylamide, tert-butylacrylamide, acrylic acid and the cross-linking agent, N,N-methylenebisacrylamide onto the surface using free radical polymerization (Maharjan et al., 2009). 5.4.2 Applications Ion-exchange chromatography There have been a number of studies examining the effect of pH and temperature-responsive resins on retention and separation of biomolecules by ion-exchange chromatography. Temperature-responsive poly-NIPAAmco-acrylic acid-co-N-tert-butylacrylamide silica beads with cation-exchange functionality have been synthesized and employed to separate bioactive peptides and amino acids. This temperature-responsive silica resin showed greater retention and separation of basic bioactive peptides at higher temperatures
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Novel adsorbents and approaches for nutraceutical separation 161 than at lower temperatures when using aqueous conditions (Kobayashi et al., 2002; 2003). Similarly, silica beads grafted with poly-NIPAAm-cobutylmethacrylate-co-dimethylaminopropyl acrylamide show better separation of small bioactive molecules (anti-inflammatory compounds) at 50 °C than at 10 °C when using an isocratic aqueous mobile phase (Ayano et al., 2006). Although changes in the retention and separation behavior of molecules have been demonstrated with temperature responsive silica resins, there have been a minimal number of reports showing the effective capture and release of biological molecules. Research recently published using porous cross-linked agarose resins has demonstrated the potential of temperature-responsive cation-exchange agarose-based resins to selectively capture and isolate charged cationic proteins such as lactoferrin from mixed protein systems (Maharjan et al., 2009). The temperature-responsive agarose resin had a 3-fold higher lactoferrin adsorption capacity at 50 °C than at 20 °C (Fig. 5.6), and was able to selectively adsorb lactoferrin from a mixture of lactoferrin, a-lactalbumin and b-lactoglobulin. In addition, dynamic studies indicated that approximately 50% of the adsorbed lactoferrin could be eluted from the agarose by simply reducing the temperature. The remaining lactoferrin was eluted using low concentrations of NaCl (Maharjan et al., 2009). Hydrophobic-interaction chromatography A significant amount of research has been undertaken to develop and understand temperature-responsive resins for hydrophobic-interaction chromatography. For example, temperature-responsive poly-NIPAAm-co-butyl methacrylate functionalized silica beads have been employed to separate a mixture of insulin chains A and B, and b-endorphin fragment 1–27 (Fig. 5.7). Temperature Adsorbed lactoferrin (mg mL–1)
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Fig. 5.6 Equilibrium adsorption isotherms of lactoferrin onto poly-NIPAAm-coacrylic acid-co-N-tert-butylacrylamide agarose resin at 20 °C (continuous line) and 50 °C (dashed line) (error bars represent the standard deviation from the mean) (Maharjan et al., 2009). Reproduced with permission. Copyright Elsevier (2009).
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162 Separation, extraction and concentration processes 1, 2, 3 2
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Fig. 5.7 Separation of a mixture of insulin chains A and B, and endorphin fragment 1–27 on a poly-NIPAAm-co-butyl methacrylate modified silica resin. Column temperature: (a) 0 °C, (b) 40 °C. Mobile phase: aqueous 0.17M NaCl. Flow-rate: 0.5 ml min–1. Detection: 215 nm (Kanazawa et al., 1997). Reproduced with permission. Copyright Elsevier (1997).
gradients were able to alter the hydrophobicity of the resin, resulting in temperature-modulated peptide elution from the column (Kanazawa et al., 1997). In addition, temperature-responsive poly-NIPAAm silica beads have been employed to separate steroids with an aqueous mobile phase (Kanazawa et al., 1996). The steroids were unresolved at temperatures below the LCST and well resolved above the LCST. The improved resolution at temperatures above the LCST was most probably the result of increased hydrophobic interaction of the steroids with the poly-NIPAAm stationary phase at the higher temperatures. Size-exclusion chromatography Temperature-responsive polymers have also been applied to synthesize temperature-responsive size-exclusion resins. For example, poly-NIPAAm was grafted onto porous glass beads, the poly-NIPAAm end functionalized with mercaptopropionic acid, and the resulting resin employed for gel permeation size-exclusion chromatography (Gewehr et al., 1992). The elution time of the dextrans was substantially altered between 25 °C and 32 °C, possibly owing to a change in the effective pore size of the resin via transition of the poly-NIPAAm chains from coils to globules on the surface of the pores of the glass beads. Porous polystyrene beads grafted with poly-NIPAAm have also shown temperature-responsive size-exclusion characteristics (Hosoya et al., 1994). In this example, an increase in temperature prolonged the elution time of higher molecular weight dextrans, possibly owing to collapse of the poly-NIPAAm causing the pore size to expand, and thus permitting the dextrans to penetrate the pores.
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Novel adsorbents and approaches for nutraceutical separation 163 Affinity separations Temperature-responsive polymers can also be conjugated to proteins and employed as affinity separation materials. Cycling below and above the LCST of the temperature-responsive polymer induces reversible precipitation/ solubilization of the bioconjugates in the aqueous solution, thereby allowing separation of the bioconjugate and the target molecule. Proteins such as a-chymotrypsin (Kim & Park, 1998) and a-glucosidase (Hoshino et al., 1998) were purified using this method. In addition, temperature-dependent isolation of lactate dehydrogenase from porcine muscle was achieved using a dye-affinity agarose modified by physically adsorbing the smart polymer poly-N-vinylcaprolactam onto the surface. Using this method, the lactate dehydrogenase was adsorbed to the column at 40 °C and eluted by reducing the temperature to 23 °C and adding KCl (Galaev et al., 1994). 5.4.3 Commercialization Temperature-responsive polymers can be employed to separate a wide spectrum of molecules using a range of separation mechanisms (e.g. charge, size, affinity). However, temperature-responsive resins are not currently available commercially for large-scale separation applications. With further research and development into specific applications and scale-up, temperature-responsive resins with silica or other matrices could be employed commercially to isolate hydrophobic and charged nutraceutical molecules. Such processes should require less solvent (e.g. salt, ethanol), which would reduce the cost and environmental impact of such separations, and potentially improve retention of the biological activity of such isolated molecules.
5.5 Mesoporous molecular sieves and applications in the nutraceutical industry Mesoporous molecular sieves, also termed mobile crystalline materials (MCMs), are a family of inorganic solids with regular nanostructure, large internal surface area (>1000 m2 g–1) and void volume, and narrow pore size distribution. These remarkable materials, exhibiting defined and tailored pore diameters from approximately 2 nm through to 50 nm, offer new possibilities for designing highly effective adsorption media based upon their pore size and large internal surface area. Mesoporous molecular sieves, notably the MCM-4X family, were first reported in the early 1990s (Beck, 1991; Beck et al., 1992; Kresge et al., 1992), and have attracted considerable research attention since then as novel catalysts, guest supports in chemical reactions, and potential adsorbents for the isolation of biological components (Beck & Vartuli, 1996; Brady et al., 2008; Kisler et al., 2001a; 2001b; Selvam et al., 2001).
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164 Separation, extraction and concentration processes 5.5.1 Synthesis Mesoporous molecular sieves are prepared using a technique called liquidcrystal ‘templating’, illustrated schematically for MCM-41 in Fig. 5.8. In this process, mesoporous solids are synthesized by thermal treatment and ultimate calcination of aluminosilicate gels in the presence of suitable surfactants. Unlike other mesoporous inorganic materials, such as silica, the MCM family of mesoporous molecular sieves comprise regular arrays of non-intersecting uniform hexagonal channels (Fig. 5.8 and 5.9), the dimensions of which can be tailored within 1.5–50 nm through the selection of different surfactants (‘template molecules’) and reaction conditions (Gusev et al., 1996; Beck, 1991; Kresge et al., 1992). The pore size can be tailored during synthesis to accommodate biological molecules of almost any size (Fig. 5.9). Hexagonal array
Silicate
Calcination
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Fig. 5.8 Schematic illustration of the liquid-crystal ‘templating’ technique for manufacture of mesoporous molecular sieves such as MCM-41 (Kresge et al., 1992). Reproduced with permission of Macmillan Publishers Ltd: Nature. Copyright Macmillan Publishers Ltd (1992).
Fig. 5.9 Model of the molecular sieve MCM-41 illustrating methane and ethane inside one of the hexagonal pores (pore diameter ~3 nm).
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Novel adsorbents and approaches for nutraceutical separation 165 5.5.2 Advantages and disadvantages The high internal surface area, tailored pore size, and very narrow pore size distribution of these mesoporous molecular sieves are characteristics that make them ideal candidates as ‘designed media’ for size-exclusion and adsorptionbased separation of biological molecules, notably proteins, peptides and other functional ingredients (Brady et al., 2008; Kisler et al., 2001a,b). For these types of applications, the best studied mesoporous molecular sieves include MCM-41 and MCM-48 (Melo et al., 1999; Zhao et al., 1996). Although the uniform mesoporous pore size makes these materials attractive candidates for adsorption and separation of biological macromolecules, their stability in aqueous systems is limited (Kisler et al., 2001b). Such instability has been a major limitation in the widespread use of these materials for waterbased extractions, and in the separation and isolation of targeted bioactive components for use as nutraceuticals. In order to overcome this shortcoming, many researchers have modified the mesoporous materials with a hydrophobic coating to assist in stabilizing the material in water (Brady et al., 2008; Kisler et al., 2001b). Such modification of MCM-41, for example, has enhanced the utility of this mesoporous molecular sieve in biochemical applications (Brady et al., 2008; Kisler et al., 2001a; 2001b). These materials have also been studied as processing aids in water treatment and in the generation of potable water (Cooper & Burch, 1999). 5.5.3 Applications Size-based separations The utility of mesoporous molecular sieves as size-exclusion materials in the isolation and separation of proteins was first examined by Diaz and Balkus (1996) and subsequently by Washmon-Kriel et al. (2000) in studies of cytochrome C permeation and adsorption. This research showed a correlation between protein adsorption and the pore size of the molecular sieve being examined, as would be expected, and retention of biological activity, an important consideration when designing a separation protocol for isolation of bioactive components like nutraceuticals. Takahashi et al. (2000; 2001) and Vinu et al. (2004) also showed strong correlation between enzyme immobilization and pore size in a study of several macroporous molecular sieves including MCM-41. The isolation of both large and small biological molecules using MCM-41 and MCM-48 has been explored (Brady et al., 2008; Kisler et al., 2001a; 2001b; Xue et al., 2004). Trypsin, lysozyme and riboflavin have been studied as model biological solutes in order to examine the size-exclusion characteristics of these mesoporous materials and their potential as bioseparation media (Kisler et al., 2001a). Similarly, Brady et al. (2008) explored the potential of hierarchical mesoporous silica sieves in the isolation of valuable dairy bioactive components as functional food ingredients, and the feasibility of such materials as industrial-scale adsorbents.
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166 Separation, extraction and concentration processes Protein stabilization Stabilization of labile enzyme activity, both in aqueous and organic environments, may well be an important application for mesoporous molecular sieves. MCM-48 was used to immobilize penicillin G acrylase and to stabilize the activity of this enzyme (Xue et al., 2004), and was shown to be superior to MCM-41 for this application. Kim et al. (2006) provided a comprehensive review of the probable mechanisms involved in stabilization of enzymes by various nanostructures, including mesoporous molecular sieves, which include confinement, charge and hydrophobic interaction, and multipoint attachment to the protein. Wang & Caruso (2005) reported the usefulness of mesoporous silica and modified forms thereof as supports for immobilization and encapsulation of enzymes, and stabilization of biological activity, demonstrating success of the technology with catalase. Mesoporous sieves may also provide utility in expanding the application scope of enzymes for processing in organic solvents by stabilizing their catalytic activity in such environments (Takahashi et al., 2001; Yan et al., 2002). 5.5.4 Commercialization The highly regular structure, tailored and tuneable properties, uniform pore size distribution and high internal surface area of selected mesoporous sieves represent attractive properties for the design of modern separation materials and protocols based upon size-exclusion and surface-interaction chemistry. However, before their industrial potential can be realized and these characteristics exploited, the long-term stability and safety, effectiveness in processing agrifood streams, and scaleability and process economics must be established.
5.6 Peptide affinity ligands and phage display methodology and applications in the nutraceutical industry Affinity chromatography is a method for separating biochemical components, often macromolecules (proteins, peptides, nucleic acids), from complex mixtures based on a highly specific biological interaction such as that between an antigen and antibody, substrate and enzyme, or ligand and receptor. The rationale for this process was first reported in the early part of the 20th century (Starkenstein, 1910). However, it was not until the late 1960s, associated with developments in modern chromatographic resins and support materials, and chemistry/linking reactions for immobilization of functional ligands, that affinity chromatography emerged as a viable industrial technology for the isolation of valuable biological components, including nutraceuticals (Cuatrecasas & Anfinsen, 1971; Wilchek & Miron, 1999). Although affinity chromatography is a very powerful technique and has revolutionized the fields of modern biochemistry, biotechnology and medical science over the past 40 years, it is not without its challenges. These relate © Woodhead Publishing Limited, 2010
Novel adsorbents and approaches for nutraceutical separation 167 to the preparation of the affinity adsorbent (ligand and resin support) and the use of the adsorbent in processing, and include considerations of selectivity, functionality, stability, efficiency, capacity, repeatability, and productivity. The choice of functional ligand to be immobilized and to effect the affinity separation is vast, and is dictated by the characteristics of the molecule targeted for isolation. Such functional ligands include proteins, enzymes, peptides, nucleic acids, sugars, vitamins, lipids, synthetic organic molecules (e.g. dyes), and metal ions. Once the ligand is selected, immobilization chemistry, maintenance of biological functionality, stability, avoidance of steric effects associated with the resin support, and effectiveness of the adsorbent under processing conditions, among others, must be considered (Wilchek & Miron, 1999). In the design of modern affinity adsorbents, the selection of ligands with broad applicability that meet most of these requirements is important in addressing the cost-effectiveness consideration critical for industrial processes. In this regard, peptides offer an attractive option (Amatschek et al., 2000; Pflegerl et al., 2002). 5.6.1 Advantages and disadvantages The main advantage of small peptides is their high affinity and selectivity. In addition, they are usually stable to mild chemical, physical and biological conditions, and not prone to degradation by protease enzymes found in most agri-food streams. Peptides are also amenable to a variety of linkage chemistries that provide for a very stable covalent bond between peptide and resin support. Tailored peptides of almost any composition and sequence can be readily synthesized through modern solid-phase techniques. Parallel and combinatorial approaches to directed peptide synthesis can readily result in large peptide libraries which can be rapidly screened for their affinity to a target compound (Pflegerl et al., 2002). Although peptide affinity resins have high affinity and selectivity, they can have relatively low binding capacity, increasing the overall cost of nutraceutical manufacture and the cost of the final product. In addition, manufacturing specific peptides on a large scale can be time consuming and expensive. Research groups are currently working on methodologies to improve the binding capacity of peptide adsorbents and significant competition in the field of peptide production is rapidly improving manufacturing technologies and reducing costs. There are also concerns regarding the toxicity and biocompatibility of novel affinity peptides when employed to produce products for human consumption. Before peptide affinity ligands can be employed in the food industry, detailed studies to ensure that they are stable during processing and non-toxic will be required. 5.6.2 Peptide identification A technique developed recently that will perhaps revolutionize the development and selection of peptide ligands for affinity adsorbtion and other applications
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168 Separation, extraction and concentration processes is called phage display, based on a concept reported in the mid-1980s (Smith, 1985). Phage display is a powerful technique for the study of protein–protein, protein–peptide, and protein–nucleic acid interactions. This approach uses bacteriophage vectors to express a protein or peptide on the surface of the virus particle coded by genetic material residing on the inside of the particle. Phage display thus creates a connection between genotype and phenotype, and facilitates the creation of a library of variants of a peptide (or protein) expressed on the outside of the virus, and allows for rapid in vitro screening of these variants based upon their binding affinity to a target molecule (e.g. enzyme, antibody, receptor, cell structural protein). This in vitro selection process is called panning, and allows for the high-throughput screening of protein interactions (Kay et al., 2001; Koivunen et al., 1999). In its simplest form, panning is carried out by incubating a library of phage-expressed peptides with the target molecule, usually in a 96-well plate, washing away any unbound phage, and then eluting the specifically bound phage. The eluted phage is then taken through additional binding and amplification cycles to enrich the pool in favor of the specific binding peptide sequence (Kay et al., 2001). In this way, phage display/panning allows for the rapid discovery of useful peptide ligands with specific affinity for the target protein or other molecules (nutraceuticals), and thus the development of effective peptide affinity adsorbents for research and industrial applications. The phage display/panning technique has been widely used in drug discovery, notably for the identification of promising new ligands (e.g. enzyme inhibitors, and receptor agonists and antagonists) to target proteins. Perhaps the most successful commercial outcome from phage display/panning technology has been the therapeutic antibody drug adalimumab (Bain & Brazil, 2003; Scheinfeld, 2003). This antibody, developed by Cambridge Antibody Technology (Cambridge, UK), targets the inflammatory cytokine TNF-a and is marketed commercially by Abbott Laboratories as HUMIRA (‘human monoclonal antibody in rheumatoid arthritis’) (Scheinfeld, 2003). 5.6.3 Commercialization The use of phage display/panning in the discovery of specific peptide– biomolecule interactions that may lead to effective affinity adsorbents was simplified through the development and sale of commercial ‘kits’ containing pre-made random peptide libraries, as well as a cloning vector for construction of custom libraries. Such ‘kits’ and related reagents are available from a number of commercial biotechnology/analytical companies, including for example New England BioLabs (Ipswich, USA), V-Biolabs (Sawston, UK), Mimotopes (Melbourne, Australia), and Cambridge Antibody Technology (Cambridge, UK) (now part of AstraZeneca). Development and refinement of phage display technology by the pharmaceutical industry provides an example of a biotechnology that can be applied in food processing applications and nutraceutical isolation. Peptide
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Novel adsorbents and approaches for nutraceutical separation 169 affinity technologies are not yet commercially available for the isolation of nutraceuticals at commercial scale. However, the phage display technology represents a robust tool for research into novel peptide affinity adsorbents. With further research and development to optimize capacity and manufacturing procedures, peptide affinity resins aimed at cost-effective isolation and separation of nutraceuticals should become available in the near future.
5.7 Membrane adsorbers, membrane chromatography and applications in the nutraceutical industry A recent development in the arena of chromatographic adsorbents is that of membrane adsorbers aimed at improving the efficiency and simplicity of the chromatographic process by allowing the targeted separation to occur on and within a macroporous membrane support (see Charcosset, 1998; Klein, 2000; Przybycien et al., 2004 for reviews of alternative bioseparation processes including the use of membrane adsorbers). This process has also been termed membrane chromatography (Charcosset, 1998; Thömmes and Kula, 1995). Membrane adsorbers are porous membranes with functional ligands (such as ion-exchange or affinity groups) covalently attached to the membrane (Demmer et al., 1989) or immobilized within the membrane structure. 5.7.1 Advantages and disadvantages Membrane adsorbers have several potential advantages over traditional particle matrices and adsorbents (Charcosset, 1998). Ion-exchange and other functionalized membranes avoid problems with swelling or packing and can easily be scaled up. Membrane adsorbers have low compressibility characteristics and also tend to demonstrate lower back pressures than traditional chromatographic resin materials thus allowing for the rapid processing of large volumes of fluid. Large targeted molecules, such as proteins and large peptides, are convected through the membrane, rather than having to diffuse into a resin bead before adsorbing to the functional ligand, and thus diffusion limitations are minimized. Overall mass transfer is improved, leading to shorter cycle times and higher overall throughput (Brandt et al., 1988). These properties facilitate the use of processing, maintenance and cleaning protocols more diverse and adaptable to existing procedures in the food processing industry. Based on the advantages of membrane adsorbers over traditional chromatographic resins (Charcosset, 1998), these adsorbents, once scaled to industrial size, would appear to be best suited for the separation and isolation of targeted components, notably proteins and peptides, found in low concentrations in large volumes of liquid (Gebauer et al., 1997). Membrane adsorbers have similar disadvantages to traditional membrane technologies. For example, when processing protein
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170 Separation, extraction and concentration processes containing fluids, a secondary boundary layer can form, reducing flux rates and increasing membrane cleaning requirements. In addition, as membrane adsorbers are made from organic materials, they are sensitive to the elevated temperatures and cleaning regimes often used in the food industry. These membranes must therefore be cleaned using mild conditions and replaced when their adsorption capacity or flux rates fall outside critical process specification. As with traditional membranes, large pressure drops across the membrane should also be avoided as they can cause deformation and irreversible changes in pore-size characteristics. 5.7.2 Functionalization The functional ligand chemistry available with membrane adsorbers is almost endless and is similar to that for traditional chromatographic resins, including strong and weak anion and cation exchange, affinity interaction, reversed phase, and hydrophobic interaction. Selection of the appropriate ligand chemistry, and membrane support material and configuration, will depend upon the specific targeted component of interest and separation strategy, source and ultimate end-use application of the isolate. 5.7.3 Configurations Improvements in membrane materials, equipment for conducting membrane chromatography, and the means by which the targeted fluid is presented to the membrane give the user choice in designing the best combination of membrane adsorber, device, and geometric configuration to achieve the most cost-effective separation of the targeted compound. Macroporous functionalized membranes are available as single or stacked flat sheets, hollow fibers, spiral-wound sheets, dead-end/single-use filters, and cassettes from several established commercial manufacturers (e.g. Pall, Sartorius) and some new entrants in the field (e.g. Mosaic Systems, Natrix Separations). These configurations allow for simple dead-end filtration applications, but also the more commercially relevant cross-flow processing for the treatment of large volumes of liquid in a short timeframe. The benefits of membrane adsorbers are most apparent in flow-through applications because capacity constraints in retention mode, particularly at high loading rates, make resin chromatography a more attractive option for capture steps. 5.7.4 Applications Ion-exchange separations Ion-exchange (anion and cation) functionalized membrane adsorbers have been the most extensively researched as cost-effective alternatives to traditional resin chromatography for the isolation of valuable minor protein and peptide constituents from agri-food streams, notably dairy whey (Adisaputro et al.,
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Novel adsorbents and approaches for nutraceutical separation 171 1996; Mitchell et al., 1994; Yang et al., 1999; Zietlow & Etzel, 1995). Attention has been devoted to the use of strongly acidic membrane adsorbers in the isolation of lactoferrin and lactoperoxidase from cheese whey (Chiu & Etzel, 1997; Mitchell et al., 1994), although some attention has also been paid to the growth factor constituents (Mitchell et al., 1994). Glycosylated and non-glycosylated forms of the glycomacropeptide have been isolated from whey using anion-exchange membrane chromatography (Kreub & Kulozik, 2009; Kreub et al., 2008). In the former study, the authors describe a direct-capture method for the isolation of glycosylated glycomacropeptide at pilot scale using a membrane adsorption process. Membrane ion-exchange chromatography has also been described for the isolation of bovine serum albumin (He & Ulbricht, 2008; Sarfert & Etzel, 1997), soybean trypsin inhibitor (He & Ulbricht, 2008; Josic & Strancar, 1999), patatin from potato (Alt et al., 2004), a-lactalbumin (Yang et al., 2002), a large protein, thyroglobulin (Yang et al., 2002), and in the large-scale production of antibodies (Zhou & Tressel, 2005; 2006; Zhou et al., 2006). Affinity separations Integration of membrane processing and affinity chromatography by way of affinity membrane adsorbers provides a number of advantages over traditional affinity chromatography with resin packed columns, notably in respect to time and recovery of bioactivity (Brandt et al., 1988; Klein, 2000; Zou et al., 2001). Such membrane adsorbents have been applied to the isolation of pepsin and chymosin using an immobilized pepstatin A membrane (Suen & Etzel, 1994), and to the separation of specific antibodies using membraneimmobilized bovine serum albumin, soybean trypsin inhibitor, Protein G, and immunoglobulin G (Kochan et al., 1996; Ostryanina et al., 2002). Affinity membranes can be dominated by slow sorption kinetics thus potentially impacting the performance of the separation (Suen & Etzel, 1994). Many of the protein and peptide examples shown have the potential to serve as valuable nutraceuticals in functional food applications (Etzel, 2004; McIntosh et al., 1998), and as natural antibacterial and bioactive agents and thus enter areas of both human and veterinary medicine, and biotechnology (Dionysius et al., 1993; Perraudin, 1991; Pouliot & Gauthier, 2006; Regester & Belford, 1999; Smithers, 2004; 2008). Micro-organism removal Membrane adsorbers provide an attractive alternative to traditional resinbased chromatography columns used to remove trace impurities and microbial contaminants (e.g. virus particles) in downstream processing applications, notably those requiring good manufacturing practice (GMP) compliance (Etzel & Riordan, 2006; Zhou et al., 2008). These membranes have been reported to reproducibly achieve greater than 4-log removal of mammalian viruses and greater than 3-log removal of endotoxin and contaminant DNA (Zhou et al., 2008).
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172 Separation, extraction and concentration processes 5.7.5 Commercialization Pall Filtration (Port Washington, USA) and Sartorius (Goettingen, Germany) are perhaps the best known manufacturers of membrane adsorbers for research and small industrial use. These manufacturers have targeted agrifood, biotechnology and biopharmaceutical processing applications for their membrane adsorber products, primarily the ion-exchange offerings. There are a number of citations in the scientific literature on the effective use of membrane ion exchangers in the isolation of biological molecules (He & Ulbricht, 2008; Mitchell et al., 1994; Przybycien et al., 2004; Yang et al., 1999). Several new manufacturers have introduced further innovations into the field that improve the inherent efficiency and simplicity of membrane adsorbers. For example, Mosaic Systems (Ba Breda, The Netherlands) is a biotechnology enterprise with a specialization in commercial-scale separation and isolation of components for the food ingredient and biopharmaceutical industries. This company has developed a novel mixed-matrix membrane adsorber that comprises functionalized resin beads (5–10 mm) immobilized in a macroporous membrane. The company claims that the embedded resin beads retain full chromatographic functionality resulting in the products’ selectivity and binding capacity, but that the macroporous membrane decouples pressure drop from the resin particle size leading to enhanced hydrodynamic performance. This mixed-matrix system leads to improved throughput and productivity compared with an equivalent traditional chromatographic system, minimal fouling, reduced water and buffer use, and retention of efficiency and specificity through multiple adsorption cycles. Mosaic Systems claims their mixed-matrix adsorbers have been proven in trials with customers for the manufacture of food ingredients, biopharmaceuticals, and the isolation of valuable components from dairy streams. Natrix Separations (Burlington, Canada) (formerly Nysa Technologies) has developed a patented polymeric hydrogel technology that combines the high binding capacity, selectivity and specificity associated with traditional chromatographic resins with the high throughput and ease of use of macroporous membranes (Childs et al., 2008). The membrane adsorber consists of a polymeric hydrogel formed within a flexible porous support matrix. The matrix provides mechanical strength, whereas the hydrogel characteristics determine the separation chemistry of the product. The hydrogel polymer provides high binding site density, a large surface area for binding, and rapid mass transfer supporting high flow rates while providing highly efficient capture of the target molecule (Childs et al., 2008).
5.8 Conclusions and sources of further information and advice There appear to be many advantages to the application of ‘next-generation’ separation materials, such as those covered in this chapter, in the isolation © Woodhead Publishing Limited, 2010
Novel adsorbents and approaches for nutraceutical separation 173 of nutraceuticals. These advantages include improved specificity, selectivity, simplicity, robustness, productivity and reduced environmental impact. However, several of these materials do remain on or over the horizon from an industrial-scale perspective. Further research is needed to establish: (i) large-scale manufacturing procedures for a number of these novel adsorbents; (ii) specific scaled applications in nutraceutical manufacture; (iii) their processing advantages over traditional chromatographic materials and devices; and (iv) perhaps of most importance from an industrial perspective, process economics and commercial feasibility. The reader is directed to the following websites for further information about the novel adsorbents and approaches covered in this chapter. ∑ ∑ ∑ ∑ ∑ ∑
Molecular imprinted polymers: www.MIPtechnologies.com www.molecular-imprinting.com http://www.sigmaaldrich.com/analytical-chromatography/samplepreparation/spe/supelmip.html Organic monoliths: www.bio-labs.com www.BIAtechnologies.com www.dionex.com Stimuli-responsive polymers: http://macromolecules.case.edu/research_stimuli-responsive.htm Mesoporous molecular sieves: www.mrs.org Peptide affinity ligands and phage display methodology: www.neb.com www.vbiolabs.com www.mimotopes.com www.cambridgeantibody.com Membrane adsorbers and membrane chromatography: www.pall.com www.sartorius.com www.mosaicsystems.nl www.natrixseparations.com
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180 Separation, extraction and concentration processes
6 Advances in the effective application of membrane technologies in the food industry M. Pinelo, G. Jonsson and A. S. Meyer, Technical University of Denmark, Denmark Abstract: This chapter focuses on the recent advances in the use of membrane technology for efficient separation and concentration of solutes in the dairy and fruit juice industry, as well as advances in the purification of bioactive compounds to be used as food additives. The importance of fouling reduction is emphasized because this is necessary for membrane processes to become economically feasible. Key words: membranes, dairy industry, fruit juice, fouling, bioactive compounds, membrane bioreactors.
6.1 Introduction Membrane separation technologies have attracted much attention in the food processing industries over recent decades. The main reason for the interest is that membranes allow efficient separation and concentration of solutes without changing phase, maintaining worthy chemical and physical properties of food components and systems with particular economical relevance. This chapter primarily focuses on the use of membrane technology in the dairy and fruit juice industries. The incorporation of new, advanced membrane filtration techniques generally occurred first in the dairy industry, and was only later incorporated into other food manufacturing processes. Fouling is still the main challenge to overcome before use of membrane technologies becomes economically feasible in industrial processes. Various techniques to reduce fouling, related to the dairy industry, are discussed in the first section of the chapter. In the fruit juice industry, membrane technology is mainly used for concentration and clarification purposes. The main challenge is to
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Advances in the effective application of membrane technologies 181 reduce the volume of juice maintaining the aroma compounds, which are otherwise lost when concentration methods such as evaporation are employed. The most recent advances in reverse osmosis and membrane distillation are reviewed, as these techniques are the most promising ones to maintain aroma compounds and juice quality. Wastewater treatment in food processing is also discussed in the chapter, as regulation regarding purified water quality and landfill management are becoming more and more restrictive. We pay special attention to wastewater treatment by using membrane bioreactors (MBR), particularly anaerobic MBR, on which research has focused on the recent years. Membrane separation is also a promising down-stream technique for purification of food ingredients and additives. We present the current challenges and limitations of applying membrane separation, coupled or not with enzyme technology, to purification and fractionation of oligosaccharides, as an example of an integrated process where membrane technology can act in synergy with other unit operations and processes.
6.2 Theoretical fundamentals of membrane separation Implementation of membrane technology in the dairy and fruit juice industry is determined by the economical feasibility of the process, which is directly associated with the separation efficiency and impact of fouling. Several theories and models have been developed in order to relate separation efficiency to the separation set-up, molecular features of solute, solvent and membrane, and operational conditions during filtration. One of these models (Jonsson and Boesen, 1975; Jonsson et al., 2008) is based on the transport of solvent and solute through the membrane pores and the friction of the solute particles and the pore wall. The model relates separation efficiency to several physical parameters: bˆ cm Ê t l Jv ˆ Ê [6.1] = 1 = b + 1 – ˜ exp Á – cp 1 – R K ÁË K¯ Ë e Di ˜¯ where cm and cp are the solute concentrations on the membrane surface and at the permeate, respectively. R is the rejection factor, b is the friction coefficient (dependent on solute and pore diameter), K is the distribution factor between pore fluid and bulk solution, t is a tortuosity factor, l is the real thickness of the fouling layer, e is the fractional pore area, Jv is the volumetric flow per unit membrane area and Di is the diffusion coefficient for a component of the solute. In certain conditions, e.g. high fluxes, concentration polarization can have a particular relevance on membrane performance. There are equations to quantify the influence of concentration polarization (e.g. Jonsson, 1985):
cm – cp ˆ Ê = exp Á J v d ˜ Ë Di ¯ cb – cp
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[6.2]
182 Separation, extraction and concentration processes and
Di =
kT 3 ph di
[6.3]
where c b represents the solute concentration in the bulk solution, d the polarization layer thickness, k the Boltzmann’s constant, and h the viscosity. These models and equations can be used as a tool to determine preliminarily, within a certain error interval, the efficiency of a particular process that could eventually be implemented at a large scale.
6.3 Membrane technology in the dairy industry Use of membrane technology in the dairy industry represents ~40% of the total applications of membranes in food processes (Daufin et al., 2001). Both polymeric and ceramic membranes are used for concentration purposes in the dairy industry. Among the polymeric ones, spiral cartridges of polysulfone or polyethersulfone are usually chosen, and tubular ceramic membranes of zirconium or aluminum oxide are also widely used. Ultrafiltration is the most widely used membrane separation process in the dairy industry (D’Souza et al., 2005). Ultrafiltration is used, for example, for concentration of milk before the cheese-making process or to reduce shipping volumes as well as for whey concentration, resulting in a protein-enriched stream which is being examined for novel applications, and which may have potential antiinflammatory and anti-cancer properties. 6.3.1 Fouling as a major concern in the dairy industry Fouling is the major concern for the dairy industry nowadays. In recent years, research efforts have aimed at the development of ultrafiltration systems with enhanced life time that can efficiently reduce the concentration polarization and fouling (Nigam et al., 2008; Rice et al., 2009a). Concentration polarization is provoked by an increasing solute concentration on the membrane surface owing to the transport of solute during convective solvent flux, which is rejected and accumulated on the membrane surface. The accumulation results in a back-diffusion of solute into the solution in the retentate side, hence creating a concentration gradient. As a consequence, an initial drop of permeate flux is observed, followed by a further progressive decrease caused by fouling. During microfiltration of skim milk for removal of bacterial spores, Guerra et al. (1997) observed a peak of permeate flux at 0.25 bar of transmembrane pressure, followed by a continuous decrease of permeate flux at increasing pressures (Fig. 6.1). After reaching 1 bar, the pressure was diminished, resulting in lower permeate fluxes than the ones obtained during
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Advances in the effective application of membrane technologies 183 the increasing pressure phase. This confirmed the influence of fouling on the permeate flux. Although theoretically the flux stabilizes when the steady-state is reached, in practice, the flux continues to decrease slowly with continued filtration. Also, fouling reduces dramatically the useful life of membranes (Rice et al., 2009b). The cleaning treatment is a key factor influencing the whole plant productivity. 6.3.2 Cleaning of membranes in the dairy industry D’Souza et al. (2005) made a complete study about the causes of membrane fouling as well as the main cleaning processes used in the dairy industry. They classified the cleaning agents into six different categories: (1) alkalis, which hydrolyze proteins and carbohydrates; (2) acids, for dissolving salts and oxides; (3) enzymes, particularly proteases and lipases; (4) surfactants, which decrease surface tension and charge; (5) sequestrants, for removal of minerals; and (6) disinfectants, since one of the main targets of the cleaning process is also the reduction of microbial load. A cleaning process generally involves three steps: product removal from the membrane and water rinsing; use of a single or a combination of appropriate cleaning agents followed by water rising; and disinfection (Trägårdh, 1989). In most cases, an acid–alkali treatment is used, the concentration of each cleaning agent being a critical parameter for each cleaning situation (Tohammadi et al., 275 250 225 Flux (L h–1 m–2)
200 175 150 125 100 75 50 25 0 0.0
0.1
0.2
0.3
0.4 0.5 0.6 0.7 0.8 Transmembrane pressure (bar)
0.9
1.0
1.1
Fig. 6.1 Permeate flux versus transmembrane pressure during microfiltration of skim milk for removal of bacterial spores using a ceramic membrane.The arrows show the direction in which the transmembrane pressure was changed. Adapted from Guerra et al. (1997).
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184 Separation, extraction and concentration processes 2002). Moreover, temperature and mechanical aspects i.e. flow rate of the cleaning solution and applied pressure during the cleaning process have to be carefully selected for each particular situation (Bartlett et al., 1995; Bird et al., 1995). The election of a proper cleaning method is also an important factor to consider in nanofiltration and reverse osmosis processes, which also have a major presence in the dairy industry. Nanofiltration is currently used, for example, to remove salt from whey as well as for the recovery of lactose from the permeate of the ultrafiltration step. Reverse osmosis has proved to be useful for whey concentration and for obtaining concentrated lactose by removing salt and minerals. 6.3.3 Recent advances for fouling reduction in the dairy industry In recent years, patents and publications on membrane technology applied to the dairy industry have focused, almost exclusively, on fouling reduction by (a) proposing new cleaning processes, (b) by optimizing the existing ones or (c) by coupling devices to the membrane set-up particularly designed to reduce the fouling, e.g. rotation/vibration and electric fields (see Section 6.7). Some have reported the advantage of using rotating or vibrating systems to enhance permeate flux and membrane selectivity (Frappart et al., 2006; Genkin et al., 2006; Jaffrin, 2008), providing promising results ready to be extrapolated to large-scale processes. Al-Akoum (2006) found a strong correlation between the permeate flux and the vibration speed in ultrafiltration of soy milk at 25 °C. The permeate flux increased ~10 times when using disks with 6-mm vanes at 11,000 rpm compared with smooth disks at 2500 rpm. Genkin et al. (2006) doubled the critical flux when using transverse vibration on submerged hollow fiber membranes, and with combined axial and transverse vibration, a five-fold increase of critical flux was observed. Frappart et al. (2006) increased the permeate flux from 130 to 230 L h–1 m–2 during recovery of lactose from diluted milk when using a rotating disk with vanes at 2000 rpm compared with a smooth disk at 1000 rpm. The same authors compared the performance of a vibrating system to the rotating disk, concluding that the rotating system is more efficient in reducing the fouling, owing to its higher shear rate. A recent patent (Makardij-Tossonian, 2009) proposes the use of an enzyme solution propelled by injecting gas to synergistically combine the effect of the enzyme solution on the breakdown of the particles causing fouling with the mechanical de-fouling effect promoted by the gas. The patent was especially developed for dairy food and beverage processing. Other inventions are more focused on the use of cleaning solutions than on the use of mechanical systems to prevent fouling. Flemming and Skou (2006) patented a system to reuse the cleaning solution by filtering it through the membrane, providing good results for ultrafiltration of dairy products. Omprakash et al. (2008) found that a 0.2% caustic solution provided the best performance for cleaning of fouled membranes used for concentration of whey protein. Measurement of fouling is another critical factor which
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Advances in the effective application of membrane technologies 185 has also been the subject of some patents. Mickols et al. (2008) conceived a system for on-line measurement of fouling in spiral wound membranes with particular potential for the dairy industry. It is based on the difference between acoustic impedances on the outer surface for a cylindrical wall of a permeate collection tube. Other systems applied other physical principles to measure the degree and nature of fouling, but the high costs associated with their implementation on a large scale is still hampering utilization at an industrial level. It can be concluded that the implementation and degree of profusion of membrane technology into the dairy industry over the coming years will be dependent on the economic impact that the several presented strategies for fouling reduction can have on the overall economy of the process. The most promising strategies are: (1) mechanical removal of fouling by using, for example, vibrating systems; (2) optimization of the conditions of use of specific chemical agents; and (3) use of enzymes able to breakdown the structure of the particular compounds causing the fouling, e.g. proteases for whey solutions (Nigam et al., 2008).
6.4 Membrane technology in the fruit juice industry Membrane technology is mainly used for clarification and for concentration of fruit juices, replacing the traditional evaporation (Table 6.1). There are several advantages of membrane processes over evaporation in fruit juices manufacture: (1) a higher quality of the final juice because of the reduction in the loss of aroma and other compounds with nutritional value; (2) lower energy consumption; and (3) versatility of the equipment which can be used to treat different products (Jiao et al., 2004). Use of traditional vacuum evaporation for concentration of fruit juices usually results in generation of off-flavours, the loss of aroma compounds and, in some cases, color degradation. Despite the drawbacks mentioned, vacuum evaporation is still the most widely used technique for concentration of juices in industry (GEA, 2009). This is probably ascribable to the fact that evaporation permits higher concentration efficiency than membrane treatments, which also in general, have, higher operational costs. 6.4.1 Reverse osmosis Amongst the membrane separation techniques used for juice concentration, reverse osmosis is the one of major interest for industry. The main advantage
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186 Separation, extraction and concentration processes Table 6.1 Some examples of utilization of membrane technology for concentration and clarification of fruit juices Application
Type of filtration
Conditions
Reference
Concentration of apple juice
Reverse osmosis
Spiral wound cellulose acetate and polyamide membranes 21.1–26.7 °C
Chua et al., 2007
Concentration of apple juice
Reverse osmosis and ultrafiltration
Plate and frame cellulose acetate membrane 20 °C
Sheu and Wiley, 1983
Clarification of cactus pear juice
Ultrafiltration
Hollow fiber polysulfone membrane 25 °C
Cassano et al., 2007a
Concentration of cactus pear juice
Osmotic distillation
Microporous polypropylene Cassano et al., hollow fiber 28 °C 2007a
Clarification of kiwifruit juice
Ultrafiltration
Hollow fiber polysulfone membrane 20–30 °C
Cassano et al., 2007b
Concentration of pineapple juice
Osmotic evaporation
Polytetrafluoroethylene membrane 25 °C
Hongvaleerat et al., 2008
Concentration of Reverse osmosis blackcurrant juice
Tubular polyethersulfone membrane 25 °C
Banvolgyi et al., 2009
Concentration of noni juice
Osmotic distillation
Hollow fiber polypropylene Valdes et al., 30 °C 2009
Concentration of Membrane blackcurrant juice distillation
Hollow fiber polypropylene Kozak et al., membrane 15 and 19 °C 2009
Recovery of the Vacuum main pear aroma membrane compound (ethyl distillation 2,4-decadienoate)
Hollow fiber module of polypropylene 25 °C
Diban et al., 2009
is that the process is carried out at low temperatures, which results in higher retention of juice components and reduction in the consumption of energy. The efficiency of the process is highly determined by the operational variables and type of membrane. Polyamide and cellulose acetate are the most common materials used for reverse osmosis of fruit juices. Polyamide membranes used for concentration of apple juice have been found to be more resistant, and have provided higher retention of flavors and higher flux than the cellulose acetate ones, and similar results were obtained for other fruit juices (Chua et al., 1987; Sheu et al., 1983; Palmieri et al., 1990). A different configuration also seems to play a main role in the performance of reverse osmosis. The plate and frame configuration have, in general, a higher retention capacity for flavor aromas than the spiral wound, owing to the lower membrane package density and area (Chou et al., 1991). Increases in transmembrane pressure and decreasing temperatures favored the retention of flavor compounds such as ethyl-2-methyl butanoate, hexanol and hexanal in concentration of apple juice (Alvarez et al., 1998; Chou et al., 1991). There are no results
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Advances in the effective application of membrane technologies 187 of advances or a significant research focus on reverse osmosis applied to concentration of fruit juices in recent years. This is probably ascribable to the limited capacity of reverse osmosis to achieve high concentration levels. Jiao et al. (2004) set 25–30 °Brix as the maximum achievable concentration of fruit juice when reverse osmosis is used, the low efficiency ascribable to the osmotic pressure, versus ~80 °Brix that can be reached by multistep vacuum evaporation. This suggests the use of reverse osmosis as a pre-step in the concentration process, which is coupled, for instance, to a subsequent evaporation. This is already being implemented in commercial plants (Gadea, 1987), being energy-saving and resulting in a higher production capacity. 6.4.2 Membrane distillation Membrane distillation can be defined as the transfer of a target solute between two solutions subjected to different temperatures and separated by a hydrophobic membrane. The temperature gradient creates a vapor pressure difference between the two interfaces, resulting in a water flux from the high-temperature to the low-temperature side. The volatile compounds evaporate and diffuse and/or convect across the membrane from the feed side, and are condensed in the other side (Khayet et al., 2002; Lawson and Lloyd, 1997). The maximum achievable concentration when using membrane distillation was reported to be 60–70 °Brix, which is very close to the one reachable by traditional evaporation (Jiao et al., 2004). The selectivity of the membrane for each of the components in the feed solution is determined by the vapor–liquid equilibrium. Therefore, the highest permeation rate will correspond to the component with the highest partial pressure (Mulder, 1996; Bandini and Sarti, 1999). The pore size of the membrane used for membrane distillation are usually between 0.2 and 1.0 mm and hydrophobic materials such as polyvinyldifluoride and polytetrafluoroethylene are the most used polymers (Mulder, 1996). The material of the membrane must be carefully selected. If the affinity between the material of the membrane and the properties of the feed solution is too high, an undesired wetting of the membrane can occur (Mengual et al., 2004; Mulder, 1996). As expected, the flux rate across the membrane will increase with higher porosities and thinner membranes. A higher temperature difference between juice and water results in a flux increase, however, temperature cannot be increased indefinitely, as some of the solute components of fruit juice are thermosensitive. As expected, permeate flux only slightly decreases with increasing juice feed concentration compared with reverse osmosis and, in general, flux is higher than for reverse osmosis at high concentration ratios (Calabro et al., 1994). Depending on the method by which the vapor is recovered from the membrane pores, the membrane distillation process can be performed by using several configurations (Fig. 6.2): (a) direct-contact membrane distillation (DCMD) uses an aqueous solution for vapor recovery; © Woodhead Publishing Limited, 2010
188 Separation, extraction and concentration processes Membrane
Aqueous solution
Flow Aqueous solution
Sweep gas Vapor
Vacuum or Air gap
(a) Vacuum Aqueous solution
Aqueous solution
Sweeping gas
(b)
(c) Boundary layers
Cfeed
Cfm
Cpermeate Cpm
Tfeed
Tfm Tpm Tpermeate
Membrane (d)
Fig. 6.2 Membrane distillation processes employed for concentration of juices: (a) cross-sectional view of a hydrophobic membrane in contact with an aqueous solution illustrating the vapor–liquid interfaces in MD; (b) configuration of SGMD; (c) configuration of VMD; (d) schematic illustration of temperature (T) and concentration (C) profiles in VMD and SGMD during the concentration process; membrane–feed boundary layer (fm), membrane–permeate boundary layer (pm).
(b) osmotic membrane distillation (OMD), an osmotic medium; (c) vacuum membrane distillation (VMD) couples to a vacuum system; (d) sweeping gas membrane distillation (SGMD) uses a sweeping gas; and
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Advances in the effective application of membrane technologies 189 (e) air-gap membrane distillation (AGMD), a stagnant air gap plus a cold plate. Of these, DCMD has probably been the most studied configuration. About 5 years ago, a very limited number of studies on application of membrane distillation for juice concentration or aroma recovery were available (Calabro et al., 1994; Khayet et al., 2002; Lagana et al., 2000). However, the number of studies on this topic has dramatically increased recently. Calabro et al. (1994) succeed in concentrating orange juice by membrane distillation with total retention of sugars and organic acids, but a considerable degradation of vitamin C caused by oxidation and high temperature. Temperature seems to be the most important factor for process improvement in membrane distillation. Bui et al. (2004) reported a flux increase of 3–4 times when feed temperature was 15 degrees higher in glucose solutions of 30 to 60%. A reduction of the flux is normally caused by viscosity increase when the concentration is higher than 50%. In all the studies, however, membrane distillation is still coupled to a prior evaporation or membrane filtration step. Kozak et al. (2006) used OMD as the final step for concentration of must and apple juice in laboratory and in a large-scale installation. After reverse osmosis, grape and apple juice reached a concentration of 23 °Brix, which increased up to ~62 °Brix by using OMD at 20 °C. No advances in the use of new membrane materials or setups are being investigated for improving membrane distillation performance. Instead, most of the studies are focused on the optimization of the operating variables for each particular fruit, in an attempt to enhance the total soluble solids concentration in the particular juice. Temperature was the major factor in the concentration of ethyl 2,4decadienoate, a pear aroma compound, over feed flow rate, feed concentration and pressure when a commercial hollow-fiber polypropylene membrane was used for VMD (Diban et al., 2009). Changes in temperature were also reported to have a higher effect than feed flux in the concentration of volatile aroma compounds from blackcurrant juice (Bagger-Jørgensen et al., 2004). The low temperature (10 °C) favored the concentration of blackcurrant aroma esters compared with 45 °C by VMD. Kozak et al. (2009) reported an increase of 80% of the permeate flux by increasing the temperature by only four degrees (from 15 to 19 °C) between both sides of the membrane, confirming the dramatic effect of this variable in blackcurrant juice concentration. In this case, a microfiltration treatment was used for pre-concentration, and the membrane distillation contribute to concentrate the juice from 22 to 58.2 °Brix. Many recent works have also measured antioxidant capacity and content of phenolic compounds, which are reported to exert beneficial health effects (Del Rio et al., 2010), in juices subjected to membrane distillation to prove the advantages of this technique over the traditional methods. The antioxidant capacity and the content of phenols remained constant after the membrane distillation treatment in cactus pear juice, noni juice and other red fruit juices (Cassano et al., 2007a; Valdés et al., 2009; Koroknai et al., 2008). The practically non-existent patents on the application of membrane © Woodhead Publishing Limited, 2010
190 Separation, extraction and concentration processes distillation for concentration of fruit juices suggest that interest in this topic is still confined to academic circles. However, the higher concentration degree attainable by using membrane distillation compared with reverse osmosis suggests a major use of this technique for the future. Integrated processes As stated above, reverse osmosis and membrane distillation processes are commonly coupled to pre-evaporation or, to a micro/ultrafiltration step at industrial level. Micro- or ultrafiltration has been proved to be useful not only for clarification purposes, but also to reduce the viscosity of the juice by removing solids and pectin, easing the posterior membrane processes by increasing the permeate flux (Jiao et al., 2004). Johnson et al. (1996) developed a method for concentrating orange juice, combining ultrafiltration for clarification with traditional evaporation, reaching a concentration higher than 80 °Brix. A more complex method has been proposed by Álvarez et al. (2002) for concentrating apple juice, including a membrane bioreactor for clarification, reverse osmosis for pre-concentration, pervaporation to recover aromas and traditional evaporation as a final step. Enzyme treatment and ultrafiltration has become the most usual clarification treatment. It has been used for concentration of blackcurrant juice coupled with reverse osmosis in an integrated process (Banvolgyi et al., 2009). In turn, membrane filtration techniques themselves are sometimes coupled to other techniques with the aim, in most instances of reducing fouling. Fernandes et al. (2007) patented a hybrid method that combines nanofiltration and electrodialysis for concentration and fouling removal of grape must, respectively.
6.5 Membrane technology for treatment of wastewater in the food industry Again, the dairy industry is the first food processing industry incorporating the new advances in membrane technology applied to wastewater treatment. It has been estimated that the dairy industry generates 0.2–11 l of waste water per liter of milk (Daufin et al., 2001). Dairy industry effluents have been reported as one of the highest volume industrial effluents (Wheatley, 1990). Carbohydrates, lipids and proteins are the most abundant organic compounds in the effluents of the dairy industry with a reported polluting charge of 0.2–2.5 g l–1 biological oxygen demand (Daufin et al., 2001; Perle et al., 1995). In general, biological processes are more efficient and costeffective than the physicochemical methods (Vidal et al., 2000). The high energy requirements associated with aerobic treatment plants have turned the research focus toward anaerobic treatments in recent years, with lower sludge production and no aeration requirements (Demirel et al., 2005). However, the membrane aerobic bioreactor is still quite common for treatment of food-processing wastewater in industry (Abdulgader et al., 2007). Membrane © Woodhead Publishing Limited, 2010
Advances in the effective application of membrane technologies 191 separation substitutes the clarification step used in traditional wastewater treatment, giving several advantages: (1) a significant reduction in the amount of sludge generated, (2) the capacity to treat a major volume of water using less space, and (3) a higher quality of the resulting water (Daufin et al., 2001). 6.5.1 Aerobic membrane bioreactor A membrane bioreactor couples the activated sludge process with membrane separation. An efficient pre-treatment, to avoid excessive fouling of the membrane must precede the biological treatment. In general, either microfiltration or ultrafiltration are used to separate the sludge from the treated water. The membrane can be either submerged or external. The most usual set-up is a submerged hollow fiber or plate membrane ultrafiltration membrane (Daufin et al., 2001; Melin et al., 2006). Once again, fouling is the major factor to assure high performance. Fouling depends, besides other general factors such as membrane type and hydrodynamic conditions, on the presence of compounds that must be produced by microbial metabolism or added to the sludge, e.g. polyelectrolytes (Melin et al., 2006). 6.5.2 Anaerobic membrane bioreactor The main advantage of anaerobic treatment compared with aerobic is the dramatic reduction in the amount of produced sludge, which is at least five times lower. In addition, the anaerobic treatment results in the production of gas that can be used to generate energy to be used elsewhere in the plant and allows a degree of degradation of the organic matter higher than 80% (ArrosAlileche et al., 2008). Some authors reported the need to operate at subcritical fluxes to assure an efficient control of fouling (Hughes and Field, 2006). The selection of the membrane is generally based on a complete retention of the biomass and very low solute rejection, critical for good performance of the membrane (Jefferson et al., 2000). Among the most usual membrane materials for membrane bioreactors, polypropylene and polyethylene are being industrially used (Kubota and Mitsubishi). Because fouling is a major factor in these systems, cleaning is a critical step in these operations and it can be done by chemical or mechanical treatments. Back-flushing has been presented as one of the most innovative and efficient mechanical techniques to remove fouling (see Section 6.7).
6.6 New applications of membrane technology for the food industry: concentration and fractionation of saccharides Production of nutri-functional compounds, highly claimed by consumers over the past decade for being able to promote healthy properties of food, has © Woodhead Publishing Limited, 2010
192 Separation, extraction and concentration processes promoted the development of new techniques for production of high-purity compounds, to be used as supplements or additives in food. Prebiotics are a good example. In recent years, many research efforts have targeted the production of oligosaccharide prebiotics for elaboration of products, e.g. beneficial-health supplements or human-like infant milk formulas (Bode, 2009). In most instances, prebiotics are produced by enzymatic break down of a polysaccharide of vegetal origin, which produces an oligosaccharide with bioactive properties (Beine et al., 2008). Membrane technology is the most feasible strategy for purification in industrial manufacture of enzymatically produced oligosaccharides (Table 6.2). However, the lack of systematic understanding of how properties of the membrane and the oligosaccharide affect membrane filtration has limited to a large extent the industrial development of processes for production of prebiotics (Pinelo et al., 2009). Some of the factors related to design of the most advantageous set-up are described in the next section and the influence of certain operational variables which have to be carefully considered to eventually exploit the process at an industrial level is explored. 6.6.1 Design of separation set-up In most small-scale studies, membrane filtration was performed in a batch dead-end system, which is the simplest set-up to operate (Pinelo et al., 2009). The system consists of a tank for the feed solution, which passes across the membrane module. The permeate is recovered in a proper reservoir (Swennen et al., 2005). A magnetic stirrer is commonly used to reduce fouling and to improve the reaction rates on a laboratory scale where the enzymatic reaction and the membrane filtration occur simultaneously (Sanz et al., 2005). Industrially, further research into other types of efficient mixing is required. In the laboratory, Sarney et al. (2000) used a membrane reactor equipped with a magnetic stirrer for the recovery of oligosaccharides from milk using a combined b-galactosidase treatment with nanofiltration. However, when fouling is a major factor for productivity, a cross-flow filtration system is recommended (Ramaswami et al., 2005). In cross-flow, the feed solution flows parallel to the membrane surface, sweeping the retained particles and helping to reduce the concentration polarization. This is probably the most advisable set-up for industrial purposes. A cross-flow system was used for purification of the products resulting from autohydrolysis of rice hulk in a tubular membrane (Vegas et al., 2006). Goulas et al. (2002) reported the same efficiency for dead-end and cross-flow filtration modes in the removal of glucose from enzymatically obtained oligosaccharides, but concluded that the concentration polarization was a major factor limiting the performance of the dead-end system.
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© Woodhead Publishing Limited, 2010
Substrate
Type of separation
Commercial powder of oligosaccharides from chicory rootstock
UF for separation of larger Oligosaccarides from UF cutoff: 20 kDa, impurities and NF for removal of 3 to 10 degrees of NF cutoff: 0.5–5 kDa monosaccharides polymerization (DP)
Kamada et al. (2002)
Arabinoxylan hydrolysates Commercial saccharides
Purification by UF (5, 10 and 30 kDa) Removal of di- and monosaccharides Purification of oligosaccarides and removal of lactose by NF Simultaneous concentration and purification by 1 kDa by NF Separation of di- and monosaccharides by NF
Stirred cell at 4 bar
Swennen et al. (2005)
Stirred cell at 50 °C and 4 bar
Sanz et al. (2005)
Caprine milk Liquors from rice husk autohydrolysis Commercial mixture of galacto-oligosaccharides
Product
Arabinoxylooligosaccharides Oligosaccharides Milk oligosaccharides Xylooligosaccharides Raffinose, sucrose and fructose
Commercial mixture of Concentration and purification by Oligosaccharides of saccharides from yacon UF and NF DP from 3 to 10 rootstock Hemicellulose hydrolyzate Purification by NF Xylose stream Enzymatic hydrolytes from vegetal origin Commercial oligosaccharide mixture Liquors from almond shells autohydrolysis
Concentration by NF Purification by NF Separation of lignin-related impurities by UF
Xylooligosaccharides Fructooligosaccharides Xylooligosaccharides
Filtration conditions
Reference
Feed velocity: 80–120 mL, 20–40 bar Sarney et al. (2000) Vegas et al. (2006) TiO2/ZrO2 ceramic membrane at 30 °C and 2–14 bar Feed concentration varied between Goulas et al. (2002) 0.15–0.08 g mL–1, filtration was performed at 25–60 °C and 5–30 bar UF cutoff 10 kDa (0.52 bar) and NF Olano-Martin et al. (2001) cutoff 5 kDa (5 bar) Concentrated feed solution: 20 wt% Retentate flow: 6 L min–1, filtration performed at 40–60 °C and 20–40 bar Feed velocity: 400 mL min–1 pH 5, 48 °C and 16 bar Feed flow: 1850 L h–1, filtration performed at 45 °C and 1 MPa Cutoffs between 1 and 8 kDa Filtration performed at 25 °C and 2.6–9 bar
Machado et al. (2000)
Huang et al. (2001) Espinoza-Gomez et al. (2004), Grassin et al. (1996) Li et al. (2004)
Advances in the effective application of membrane technologies 193
Table 6.2 Some examples of the use of ultrafiltration (UF) and nanofiltration (NF) membranes for the concentration and purification of oligosaccharides of different nature
194 Separation, extraction and concentration processes 6.6.2 Operational modes The integration of enzyme reactor and membrane is normally advantageous for recovery of prebiotic oligosaccharides from vegetal matrixes, mainly because this disposition diminishes product inhibition (Pinelo et al., 2009). The operational conditions for this particular integrated system depend to a large extent on whether the system operates with immobilized or free enzymes. On some occasions, the membrane itself was used for the immobilization. Nishizawa et al. (2000) reported the chemical immobilization of b-fructofuranosidase with glutaraldehyde to the inner surface of a ceramic membrane for production of fructo-oligosaccharides. Although the use of ceramic membranes has been used for purification of oligosaccharides by ultra- and nanofiltration, it is more common to use polymeric organic membranes for oligosaccharide purification (Artug et al., 2007). Cellulosic membranes seem more suitable to this purpose, being hydrophilic, but the use of polysulfone membranes is also common (Akhtar et al., 1995). A chemical surface treatment was recommended for a polysulfone membrane, in order to make it more hydrophilic (Akhtar et al., 1995). The operating variables that had a major influence on membrane filtration of oligosaccharides were: (1) temperature, (2) pH, (3) concentration of the feed solution, and (4) pressure. Temperature It has been reported that the permeate flux almost doubled when the temperature of saccharide solutions increased from 20 to 50 °C. This is ascribable to the reduction of viscosity, which favors the pass of the flow of the feed solution across the membrane (Goulas et al., 2002; Machado et al., 2000). The favorable effect of the temperature increase on the permeate flux was confirmed by Sjöman et al. (2008), during recovery of xylose from hemicellulose hydrolyzate feeds using various polysulfone membranes. pH For certain types of membrane carrying positive or negative charged groups, e.g. polycarbonate and sulfonated polysulfone membranes, the influence of pH on the saccharide rejection can be dramatic (Lazaridou et al., 2007). When the feed solution includes some charged compounds, flux commonly increases when working close to the isoelectric point of the membrane, as expected (Childress and Elimelech, 1996). Variation of pH can be used as a strategy for more efficient separation of charged saccharides, such as pectin, whose charge density and distribution is a function of the carboxyl groups in the structure.
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Advances in the effective application of membrane technologies 195 Concentration of the feed solution Permeate flux is expected to decrease with increasing feed concentration of the solution. As a result, a decrease in the rejection, at constant pressure, is usually observed in these cases (Goulas et al., 2002). Pressure The role of pressure and the recommended levels of pressure depend on the type and properties of the solution to be filtered. Generally, the rejection increases with increasing levels of pressure until reaching a certain level. This is ascribable to the higher flux of solute across the membrane, which decreases the difference in concentration between the solutions at both sides of the membrane. When pressure continues increasing, the concentration polarization and, in turn, the concentration of solute on the surface of the membrane increases too. As a consequence, the concentration of solute on the permeate side increases and the rejection decreases. In purification of xylo-oligosaccharides from lignin impurities, Nabarlatz et al. (2007) observed a decrease in rejection when the pressure was increased from 2.6 to 9 bar. It is then advised to operate at low pressures, within a range in which the concentration polarization is not so high, in order to maximize selectivity.
6.7 Future trends Implementation of membrane technology in food industry has been hampered by the economical limitations derived from fouling over recent decades. Many studies have been devoted to develop techniques and processes to reduce the impact of fouling and it seems that progress in membrane technology will be linked to progress in fouling reduction over the next few years. In particular, there are at least three techniques that have provided promising results for their efficiency and simplicity to be incorporated in a large-scale process: (1) high frequency backflushing, (2) vibrating membrane modules, and (3) electrofiltration. Some of the benefits associated with coupling these techniques in the dairy industry have been discussed in Section 6.3 but the applications in other food processes are increasing 6.7.1 High frequency backflushing Backflushing consists of applying pressure pulses from the permeate to the retentate side, and stripping off the particles causing fouling from the membrane surface into the bulk solution at the permeate side again (Pinelo et al., 2009). Backflushing is commonly achieved by placing a valve in the © Woodhead Publishing Limited, 2010
196 Separation, extraction and concentration processes 300
Flux (L h–1 m–2)
250 With backflush
200 150 100 50 0
Without backflush 0
15
30
45
60 75 Time (min)
90
105
120
Fig. 6.3 Influence of backflushing on the permeate flux versus time applied to the microfiltration of skim milk for removal of bacterial species using a ceramic membrane. Adapted from Guerra et al. (1997).
permeate side. The efficiency of backflushing depends both on the duration of the pulse itself and time between pulses. Guerra et al. (1997) reported a considerable increase of permeate flux when backflushing was incorporated to microfiltration of skim milk, as backflushing contributed to the removal of spores from the membrane surface (Fig. 6.3). Jonsson (2008) narrowed the intervals of fractionation of dextrans in a hollow-fiber ultrafiltration system using times between pulses from 1 to 30 s and backflushing times from 0.1 to 5 s. Application of backflushing in industry is still limited but further research in the coming years will permit a more detailed analysis of the balance between costs associated with implementation and economic benefit from fouling reduction. 6.7.2 Vibrating membrane module A vibrating design consists of applying high shear rates to favor the transmission of macromolecules through the membrane (Beier et al., 2007). This design proved to be particularly efficient for nanofiltration of diluted milk and for enzyme separation applied to food processes (Beier et al., 2007; Jaffrin et al., 2004). This system has been already presented in the section dedicated to fouling reduction in the dairy industry, but it is fair to include it here as a future trend technique, as many studies are currently being conducted to optimize the mechanical characteristics of the system for a feasible high performance of membrane separation in the dairy industry.
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Advances in the effective application of membrane technologies 197 6.7.3 Electrofiltration Electrofiltration is used when the particles causing concentration polarization have a net charge and can be attracted or repelled by applying an electric field, in this instance across the membrane (Pinelo et al., 2009). This technique has been successfully employed for separation of industrial solutions of enzymes with a surface charge, resulting in a dramatic increase of the permeate flux (Enevoldsen et al., 2007). The technique also gave good results in ultrafiltration of other proteins (Oussedik et al., 2000) and shows promise for purification of potential food ingredients in which charge can be induced, e.g. modifying the number and methoxyl groups on pectin molecules.
6.8 References Abdulgader M E, Yu Q J, Williams P, Zinatizadeli A A L (2007), ‘A review of the performance of aerobic bioreactors for treatment of food processing wastewater’. Proceedings of the International Conference on Environmental management, Engineering, Planning and Economics, pp. 1131–1136. Skiathos. Akhtar S, Hawes C, Dudley L, Reed P, Strafford P (1995), ‘Coatings reduce the fouling of microfiltration membranes’, J Membr Sci 107, 209–218. Al-Akoum O, Richfield D, Jaffrin M Y, Ding L H, Swart P (2006), ‘Recovery of trypsin inhibitor and soy milk concentration by dynamic filtration’, J Membr Sci 279, 291–300. Álvarez V, Riera F A, Álvarez S, Álvarez R (1998), ‘Permeation of apple aroma compounds in reverse osmosis’, Sep Purif Technol 14, 209–220. Álvarez S, Riera F A, Álvarez R, Coca J (2002), ‘Concentration of apple juice by reverse osmosis at laboratory and pilot-plant scale’, Ind Eng Chem Res 41, 6156–6164. Arros-Alileche S, Merin U, Daufin G, Gesan-Guiziou G (2008), ‘The membrane role in an anaerobic membrane bioreactor for purification of dairy wastewaters: A numerical simulation’. Bioresour Technol 99, 8237–8244. Artug G, Roosmasari K, Richau K, Hapke A (2007), ‘A comprehensive characterization of commercial nanofiltration membranes’, Sep Sci Technol 42, 2947–2986. Bagger-Jørgensen R, Meyer A S, Varming C, Jonsson G (2004), ‘Recovery of volatile aroma compounds from black currant juice by vacuum membrane distillation’, J Food Eng 64, 23–31. Bandini S, Sarti G C (1999), ‘Heat mass transport resistances in vacuum membrane distillation per drop’, AIChE J 45, 1422–1433. Banvolgyi S, Horvath S, Stefanovits-Banyai E, Bekassy-Molnar E, Vatai G (2009), ‘Integrated membrane process for blackcurrant juice concentration’, Desalination 241, 281–287. Barlett M, Bird M R, Howell J A (1995) ‘An experimental study for the development of a qualitative membrane cleaning model’, J Membr Sci, 218, 107–116. Beier S P, Jonsson G (2007), ‘Separation of enzymes and yeast cells with a vibrating hollow fiber membrane module’ Sep Purif Technol 53, 111–118. Beine R, Moraru R, Nimtz M, Na’amnieh S, Pawlowski A, Buchholz K, Seibel J (2008), ‘Synthesis of novel fructooligosaccharides by substrate and enzyme engineering’, J Biotechnol 138, 33–41. Bird M R, Barlett M (1995), ‘CIP optimization for the food industry: Relationships between detergent concentration, temperature and cleaning time’, Trans IChemE, 73, 63–70.
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198 Separation, extraction and concentration processes Bode L (2009), ‘Human milk oligosaccharides: prebiotics and beyond’, Nutr Rev 67, 183–191. Bui V A, Nguyen M H, Muller J (2004), ‘A laboratory study on glucose concentration by osmotic distillation in hollow fiber modules’, J Food Eng 63, 237–245. Calabro V, Jiao B L, Drioli E (1994), ‘Theoretical and experimental study on membrane distillation in the concentration of orange juice’, Ind Eng Chem Res 33, 1803– 1808. Cassano A, Conidi C, Timpone R, D’Avella M, Drioli E (2007a), ‘A membrane-based process for the clarification and the concentration of the cactus pear juice’, J Food Eng 80, 914–921. Cassano A, Donato L, Drioli E (2007b), ‘Ultrafiltration of kiwifruit juice: Operating parameters, juice quality and membrane fouling’, J Food Eng 79, 613–621. Childress A E, Elimelech M (1996), ‘Effect of solution chemistry on the surface charge of polymeric reverse osmosis and nanofiltration membranes’, J Membr Sci 119, 253–268. Chou F, Wiley R C, Schlimme D V (1991), ‘Reverse osmosis and flavor retention in apple juice concentration’, J Food Sci 56, 484–487. Chua H T, Rao M A, Acree T E, Cunningham D G (1987), ‘Reverse osmosis concentration of apple juice: flux and flavor retention by cellulose acetate and polyamide membranes’, J Food Proc Eng 9, 231–245. D’Souza N M, Mawson A J (2005), ‘Membrane cleaning in the dairy industry: a review’, Crit Rev Food Sci Nutr, 45, 125–134. Daufin G, Escudier J P, Carrere H, Berot S, Fillaudeau L, Decloux M (2001), ‘Recent and emerging applications of membrane processes in the food and dairy industry’, Trans IChemE, 79, 89–102. Del Rio D, Costa L G, Lean M E J, Crozier A (2010), ‘Polyphenols and health: what compounds are involved?’, Nutr Metab Cardiovasc Dis, 20, 1–6. Demirel B, Yenigun O, Onay T T (2005), ‘Anaerobic treatment of dairy wastewaters: a review’, Process Biochem 40, 2583–2595. Diban N, Voinea O C, Urtiaga A, Ortiz G (2009), ‘Vacuum membrane distillation of the main pear aroma juice compound. Experimental study and mass transfer modelling’, J Membr Sci 326, 64–75. Enevoldsen A D, Hansen E B, Jonsson G (2007), ‘Electro-ultrafiltration of industrial enzyme solutions’, J Membr Sci 28–37. Espinoza-Gomez H, Lin S W, Rogel-Hernandez E (2004), ‘Nanofiltration membrane pore diameter determination’, Rev Soc Quim Mex 48, 15–20. Fernandes G V M, Lopes C G, Correia N (2007) ‘Concentration and rectification of grape must involves combining nanofiltration and electrodialysis of grape must in hybrid process’, World Patent WO2008051100-A2. Flemming S, Skou F (2006), ‘Membrane filtration of a product, e. g. milk, in membrane plant, involves recovering cleaning solution after multi-step cleaning of membrane system in plant; and using recovered cleaning solution for cleaning membrane system’, European Patent EP1726353. Frappart M, Akoum O, Ding L H, Jaffrin M Y (2006), ‘Treatment of dairy process waters modeled by diluted milk using dynamic nanofiltration with a rotating disk module’, J Membr Sci 282, 465–472. Gadea A (1987), ‘Reverse osmosis of orange juice’. In Proceedings of the International fruit juice congress, Orlando, USA. GEA (2009) Report on ‘Evaporation technology for the juice industry’, GEA Wiegand GmbH, www.gea-wiegand.com. Accessed on Dec 15, 2009. Genkin G, Waite A G, Fane S, Chang S (2006), ‘The effect of vibration and coagulant addition on the filtration performance of submerged hollow fiber membranes’, J Membr Sci 281, 726–734. Goulas A K, Kapasakalidid P G, Sinclair H R, Rastall R A, Grandison A S (2002),
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Advances in the effective application of membrane technologies 199 ‘Purification of oligosaccharides by nanofiltration’, J Membr Technol 209, 321– 335. Grassin C, Fauquembergue P (1996), ‘Fruit Juices’, in: T. Godfrey, S. West (Eds.), Industrial Enzymology, Macmillan Press. Guerra A, Jonsson G, Rasmussen A, Nielsen E W, Edelsten D (1997), ‘Low cross-flow velocity microfiltration of skim milk for removal of bacterial spores’, Int Dairy J 7, 849–861. Hongvaleerat C, Cabral L M C, Dornier M, Reynes M, Ningsanond S (2008), ‘Concentration of pineapple juice by osmotic evaporation’, J Food Eng 88, 548–552. Huang R Y M, Shao P, Nawawi G, Feng X, Burns C M (2001), ‘Measurements of partition, diffusion coefficients of solvents in polymer membranes using rectangular thin-channel column inverse gas chromatography’, J Membr Sci 188, 205–218. Hughes D, Field R W (2006), ‘Crossflow filtration of washed and unwashed yeast suspensions at constant shear under nominally sub-critical conditions’, J Membr Sci 280, 89–98. Jaffrin M J (2008), ‘Dynamic shear-enhanced membrane filtration: a review of rotating disks, rotating membranes and vibrating systems’, J Membr Sci, 324, 7–25. Jaffrin M Y, Ding L H, Akoum O, Brou A (2004), ‘A hydrodynamic comparison between rotating disk and vibratory dynamic filtration systems’, J Membr Sci 242, 155–167. Jefferson B, Laine A T, Judd S J, Stephenson T (2000), ‘Membrane bioreactors and their role in wastewater reuse’, Water Sci Technol 41, 197–204. Jiao B, Cassano A, Drioli E (2004), ‘Recent advances on membrane processes for the concentration of fruit juices: a review’, J Food Eng 63, 303–324. Johnson J R, Braddock R J, Chen C S (1996), ‘Flavor losses in orange juice during ultrafiltration and subsequent evaporation’, J Food Sci 61, 540–543. Jonsson G (1985), ‘Molecular weight cut-off curves for ultrafiltration membranes of varying pore sizes’, Desalination 53, 3–10. Jonsson G (2008), ‘Tuning of the cut-off curves by dynamic ultrafiltration’, Proceedings of the International Conference on Membranes and Membrane Processes: ICOM2008, Hawaii, July 12–18. Jonsson G, Boesen C E (1975), ‘Water and solute transport through cellulose acetate reverse osmosis membranes’, Desalination 17, 145–165. Kamada T, Nakajima M, Nabetani H, Iwamoto N S (2002), ‘Availability of membrane technology for purifying and concentrationg oligosaccharides’, Eur Food Res Technol 214, 435–440. Khayet M, Godino M P, Mengual J I (2002), ‘Thermal boundary layers in sweeping gas membrane distillation processes’, AIChE J 48, 1488–1497. Koroknai B, Csanadi Z, Gubicza L, Belafi-Bako (2008), ‘Preservation of antioxidant capacity and flux enhancement in concentration of red fruit juices by membrane processes’, Desalination 228, 295–301. Kozak A, Bekassy E, Vatai G (2009), ‘Production of black-currant juice concentrate by using membrane distillation’, Desalination 241, 309–314. Kozak A, Rektor A, Vatai G (2006), ‘Integrated large-scale membrane process for producing concentrated fruit juices’, Desalination 200, 540–542. Lagana F, Barbieri G, Drioli E (2000), ‘Direct contact membrane distillation: modeling and concentration experiments’, J Membr Sci 166, 1–11. Lawson K W, Lloyd D R (1997), ‘Membrane distillation’, J Membr Sci 124, 1–25. Lazaridou A, Biliaderis C G (2007), ‘Molecular aspects of cereal beta-glucan functionality: physical properties, technological applications and physiological effects’, J Cereal Sci 46, 101–118. Li W, Li J, Chen T, Chen C (2004), ‘Study on nanofiltration for purifying fructooligosaccharides II. Extended pore model’, J Membr Sci 258, 8–15. Machado D R, Hasson D, Semiat R (2000), ‘Effects of solvent properties on permeate through nanofiltration membranes. Part II: transport model’, J Membr Sci 166, 6–69.
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200 Separation, extraction and concentration processes Makardij-Tossonian A A (2009), ‘Membrane regeneration’, World Patent WO/2009/089587. Melin T, Jefferson B, Bixio D, Thoeye C, De Wilde W, De Koning J, van der Graaf J, Wintgens T (2006), ‘Membrane bioreactor technology for wastewater treatment and reuse’, Desalination 187, 271–282. Mengual J I, Khayet M, Godino M P (2004), ‘Heat and mass transfer in vacuum membrane distillation’, Int J Heat Mass Transfer 47, 865–875. Mickols W, Koreltz M S, Moll D J, Streeter D B, Mickols W, Kobeltz M S (2008), ‘Spiral wound module assembly for e.g. in-situ on-line and real-time measurement of membrane fouling in spring wound module, has membrane envelopes and acoustic transducers which are located adjacent to permeate collection tube’, World Patent WO2008103864. Mulder M (1996), Basic principles of membrane technology, Dordrecht/Boston/London: Kluwer Academic Publishers. Nabarlatz D, Torras C, Garcia-Valls R, Montane D (2007), ‘Purification of xylooligosaccharides from almond shells by ultrafiltration’, Sep Purif Technol 53, 235–243. Nigam M O, Bansal B, Chen X D (2008), ‘Fouling and cleaning of whey protein concentrate fouled ultrafiltration membranes’, Desalination 218, 313–322. Nishizawa K, Nakajima M, Nabetani H (2000), ‘A forced flow membrane reactor for transfructosylation using ceramic membrane’, Biotechnol Bioeng 68, 92–97. Olano-Martin E, Mountzouris K C, Gibson G R, Rastall R A (2001), ‘Continuous production of pectic oligosaccharides in a membrane enzyme reactor’, J Food Sci 66, 966–971. Omprakash M, Bansal B, Cheng D X (2008), ‘Fouling and cleaning of whey protein concentrate fouled ultrafiltration membranes’, Desalination 218, 313–322. Oussedik S, Belhocine D, Grib H, Loucini H, Piron D L, Nameri N (2000), ‘Enhanced ultrafiltration of bovine serum albumin with pulsed electric field and fluidized activated alumina’, Desalination 127, 59–64. Palmieri L, Dalla Rosa M, Dall’Aglio G, Carpi G (1990), ‘Production of kiwifruit concentrate by reverse osmosis process’, Acta Horticulturae 282, 435–439. Perle M, Kimchie S, Shelef G (1995), ‘Some biochemical aspects of the anaerobic degradation of dairy wastewater’, Water Res 29, 1549–1554. Pinelo M, Jonsson G, Meyer A S (2009), ‘Membrane technology for purification of enzymatically produced oligosaccharides: molecular and operational features affecting performance’, Sep Purif Technol, 70, 1–11. Ramaswami H S, Marcotte M (2005), Food processing: principles and applications, CRC Press. Rice G, Barber A, O’Connor A, Stevens G, Kentish S (2009a), ‘Fouling of nanofiltration membranes by dairy ultrafiltration permeates’, J Membr Sci 330, 117–126. Rice G, Kentish S, O’Connor A, Barber A, Pihlajamäki A, Nyström M, Stevens G (2009b), ‘Analysis of separation and fouling behaviour during nanofiltration of dairy ultrafiltration permeates’, Desalination 236, 23–29. Sanz M L, Polemis N, Morales N, Corzo A (2005), ‘In vitro investigation into the potential prebiotic activity of potential oligosaccharides’, J Agric Food Chem 53, 2914–2921. Sarney D B, Hale C, Frankel C, Vulfson E N (2000), ‘A novel approach to the recovery of biologically active oligosaccharides from milk using a combination of enzymatic treatment and nanofiltration’, Biotechnol Bioeng 69, 461–467. Sheu M J, Wiley R C (1983), ‘Preconcentration of apple juice by reverse osmosis’ J Food Sci, 48, 422–429. Sjöman E, Mänttäri M, Nyström M, Koivikko H, Heikkilä H (2008), ‘Xylose recovery by nanofiltration from different hemicellulose hydrolyzate feeds’, J Membr Sci 310, 268–277.
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Advances in the effective application of membrane technologies 201 Swennen C M, Coyrtin B, Bruggen C, Vandecasteele C, Gibson G R, Rastall R A (2005), ‘Ultrafiltration and ethanol precipitation for isolation of arabinoxylooligosaccharides with different structures’, Carbohydr Polym 62, 283–292. Tohammadi T, Madaeni S S, Moghadam M K (2002), ‘Investigation of membrane fouling’, Desalination, 153, 155–160. Trägårdh G (1989), ‘Membrane cleaning’, Desalination, 71, 325–335. Valdés H, Romero J, Saavedra A, Plaza A, Bubnovich V (2009), ‘Concentration of noni juice by means of osmotic distillation’, J Membr Sci 330, 205–213. Vegas R, Luque S, Alvarez J R, Alonso J L, Dominguez H, Parajo R C (2006), ‘Membraneassisted processing of xylooligosaccharides-containing liquors’, J Agric Food Chem 54, 5430–5436. Vidal G, Carvalho A, Mendez R, Lema J M (2000), ‘Influence of the content in fats and proteins on the anaerobic biodegradability of dairy wastewaters’, Bioresour Technol 74, 231–239. Wheatley A (1990), Anaerobic digestion: a waste treatment technology. London and New York: Elsevier Applied Science.
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7 Electrodialytic phenomena, associated electromembrane technologies and applications in the food, beverage and nutraceutical industries L. Bazinet, A. Doyen and C. Roblet, Laval University, Canada Abstract: Electrodialysis, an electrochemical separation process with charged membranes stacked to separate ionic species from aqueous solutions and uncharged components when an electrical field is applied, is providing new membrane separation processes with numerous applications in the food, nutraceutical and beverage industries. Techniques such as electrolysis with membrane, electrodialysis with ionexchange membranes, electrodialysis with bipolar membrane and electrodialysis with filtration membrane have been used for, among other applications, the coagulation of protein, the electroreduction of the medium and/or the fractionation of several food proteins. Moreover, electrodialysis with filtration membranes has been used for the recovery of molecules with bioactive properties such as antioxidant, anticancer and antihypertensive peptides. Key words: electrodialysis, electromembranes, nutraceuticals, biomolecules, purification, separation.
7.1 Introduction Electrodialysis (ED) is one of a group of membrane-based separation technologies which are finding increasing use, in agri-food industries to concentrate, purify or modify foods. In ED, an electric field provides the driving force and porous or non-porous membranes perform the separation: electrodialysis is a combined method of dialysis and electrolysis (Shaposhnik and Kesore, 1997). ED can be performed with two main cell types: electrolysis (or electro-electrodialysis) cells for oxido-reduction reactions and multimembrane cells for dilution–concentration, water dissociation and purification applications. The electrolysis cell operates with only one membrane that © Woodhead Publishing Limited, 2010
Electrodialytic phenomena and associated electromembrane technologies 203
separates two solutions circulating in each electrode compartment. This application is based on electrode redox reactions that are electrolysis-specific properties. In multi-membrane cells only the membrane transport phenomena intervenes, whereas electrochemical reactions occurring at the electrodes do not interact with the separation process. The principles of electrode and membrane reactions, as well as the technologies associated with these electrodialytic phenomena are reviewed in this chapter. The specific applications of electrodialysis and electrolysis currently used and those under development in the food, beverage and nutraceutical industries are presented.
7.2 Principles of electrodialytic phenomena and associated membrane technologies 7.2.1 Membrane phenomena An ion-exchange membrane is made of a macromolecular material (skeleton) which carries ionizable groups such as ion-exchange resins. The membrane contains fixed ions firmly attached to the skeleton and is electronically neutralized by mobile charges of the opposite sign, called counterions. Counterions, which carry current in the membrane, are positive in the case of cation-exchange membranes (CEM), and negative for anion-exchange membranes (AEM) (Bazinet, 2005). Both of these membranes are monopolar; this means that they are permeable to only one type of ion (Gardais, 1990). The perm-selectivity of the membrane is the result of an electrostatic repulsion called Donnan exclusion (Donnan, 1911). For a cation-exchange membrane (CEM), anions are repulsed from the membrane thus allowing only cations to migrate through the membrane. 7.2.2 Electrode phenomena In membrane electrolysis, membrane selectivity, explained above, and electrolysis phenomenon are both active. In electrolysis, an external potential difference is applied to the cell and chemical reactions occur at the electrode–solution interface. The faradaic reactions, caused by the passage of a current, are characterized by electron transfer at the electrode–solution interface. The oxidation A Æ Az+ + z e– induces the loss of one or more electrons and the reduction Bn+ + n e– Æ B induces the gain of one or more electrons. The anode induces oxidation reactions, and reduction reactions occur at the cathode. This transfer always occurs at the electrode surface (Brett and Oliveira-Brett, 1994; Gardais, 1990).
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204 Separation, extraction and concentration processes
7.3 Applications of electrodialytic phenomena and associated membrane technologies Only one type of membrane technology is associated with electrode reactions (electrolysis with membrane) whereas three types are associated with membrane phenomena (electrodialysis with ion-exchange membranes, electrodialysis with bipolar membrane, and, recently, electrodialysis with filtration membranes) (Fig. 7.1). 7.3.1 Electrolysis with membrane Two main applications of electrode reactions are electrochemical coagulation (EC) and electroreduction. Both methods are exploratory, but show an interesting potential for applications in the food, beverage and nutraceutical industry (Table 7.1). An example of both applications will be developed. Milk protein electrochemical coagulation Khidirov and Merzametov (1982) used EC to precipitate milk proteins in an electrolysis cell separated by a ceramic diaphragm. The milk is poured into the anodic compartment and an electrolyte or a whey solution is poured into the cathodic compartment. Proteins coagulate on the platinum electrode by forming a highly dense white coagulum. Every type of milk could be treated for protein electrochemical coagulation without the use of rennet. Both the caseins and the whey proteins coagulate at the anode. Janson and Lewis (1994) studied the possible use of electrochemical coagulation to directly separate up to 73.8% of the total cheese whey proteins. The whey was circulated in the anodic compartment of an electrolysis cell with a non-ion-selective membrane, made of cotton, for the acidification–coagulation phase (Janson et al., 1990). Afterwards the solution was separated into the coagulum and an impoverished protein solution. This solution was then circulated in the cathodic compartment of the electrolysis cell to undergo the alkalinization phase. Enhancement of lipid stability of omega-3 enriched commercial milk Consumer demand for specific nutritional qualities is encouraging the dairy industry to develop products supplemented in omega-3 fatty acids. Although these fatty acids are known to have many health benefits, they are extremely susceptible to oxidative deterioration which causes difficulties during their storage. In the study by Haratifar (2008) an electroreduction process was performed to modify the redox state of omega-3 enriched milk. A 4-V electroreduction treatment was applied for 1 h on pasteurized omega-3 enriched milk, at room temperature. The electroreduction treatment reduced the redox potential value of omega-3 enriched milk samples quickly and decreased their dissolved oxygen concentration. The electroreduced and control samples were stored at room temperature for up to 3 weeks in the presence of fluorescent light and
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(b)
2H2O +
2e– + 2H2O H+
O2 + 4H+ + 4e– OH–
Anode
H2 +
Desalted solution
AEM CEM
AEM CEM
Cl–
– Cathode
+ Anode
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2OH–
Cl
–
Cl Na
Cl–
–
Na+ Na+
+
–
Na+
Cathode n
Feed solution (c)
Anionic/acid fraction
AEM FM
Neutral fraction
FM CEM
FM
P+/– + Anode
A
–
P–
P+ A– C+ P–
P+
A
(d)
Cationic/basic fraction
–
P–
FM CEM P+/–
Cl–
A– P+ P+ C
NaOH solution
AEM CEM BPM AEM Cl–
HCl solution CEM BPM AEM H+
H+
+
C
P–
+
–
+
Cathode
Anode
OH Na+
–
Na+ Cl–
Cl
–
Na+
n Feed solution
OH– Na+
Cl–
CEM
Cl–
–
Na+
Na+ Cathode n
Salt solution
Fig. 7.1 Electrodialytic phenomena and associated membrane configurations: (a) redox reaction; (b)–(d) membrane reactions. (a) Electrolysis with membrane; (b) electrodialysis with ion-exchange membranes; (c) electrodialysis with filtration membranes; and (d) electrodialysis with bipolar membranes. AEM, anion-exchange membrane; CEM, cation-exchange membrane; BPM, bipolar membrane; FM, filtration membrane; A–, anion; C+, cation; P–, anionic peptide; P+, cationic peptide; P+/–, neutral peptide.
Electrodialytic phenomena and associated electromembrane technologies 205
(a)
206 Separation, extraction and concentration processes in the dark and the composition in fatty acids of milk samples was measured by gas chromatography with flame ionization detection. It was shown that storage under fluorescent light involved a degradation of the fatty-acids, whereas the electroreduction treatment slowed down the oxidative degradation of electroreduced samples compared with untreated milk samples. Results of this study show that the electroreduction treatment can be a potential method of enhancing the shelf-life of products containing unsaturated fatty acids. 7.3.2 Electrodialysis with ion-exchange membranes Conventional electrodialysis or electrodialysis with ion-exchange membranes consist of a series of cation- and anion-exchange membranes arranged in an alternating pattern between an anode and a cathode to form individual cells. Under the influence of an applied potential gradient between cathode and anode, the positively charged cations migrate towards the cathode and negatively charged anions towards anode (Bazinet, 2005, Bazinet and Firdaous, 2009b). The main applications of ED and dilution–concentration in the food industry consist of the demineralization of milk, sugar and whey. In the 1990s, a new application was developed for tartaric stabilisation of wine. The use of ED in membrane bioreactors, for protein separation and acid production was studied. The main applications of conventional ED are presented in Table 7.1, and the application on tartaric stabilization of wine is described in detail here. The use of ED for tartaric stabilization of wine has been approved by the Council of Europe in 1998 and is presently used worldwide on an industrial scale. Wine contains tartaric acid (H2T), a dicarboxylic acid which dissociates in the tartrate (HT–) and bitartrate (T2–) forms. During alcoholic fermentation, potassium and tartrate ions associate and precipitate to form potassium tartrate crystals (KHT) (Gonçalves et al., 2003). As a consequence, at normal storage temperatures, an undesirable KHT precipitation occurs in wine bottles. The ED process for wine tartaric stabilization was developed by Escudier et al. (1995). The principle is the same as for the conventional ED process with a diluate compartment in which the wine to be treated circulates, under nitrogen bubbling, and a concentrate compartment where salts extracted from the wine are recovered. According to the authors, the level of wine deionization, based on the decrease of the conductivity, should be 5 to 20% to obtain tartaric stabilization of wine. One of the main advantages of this technique is the preservation of the organoleptic properties of wine during the stabilization process. 7.3.3 Electrodialysis with filtration membranes Bazinet et al. (2005) stacked filtration membranes and ion-exchange membranes in a conventional electrodialysis cell. This technology, named electrodialysis
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Technology
Substrate
Target compounds Objective
Applications and stage of development
References
Dissolved oxygen
Protection of flavour after spray-drying
Food (in development)
Inoue and Kato, 2003
Fruit juices
Dissolved oxygen
Increase in shelf-life
Food (in development)
Hekal, 1983 Swanson and Sommer, 1940
Milk
W-3 fatty acids
Protection against lipid oxidation
Functional food (in development)
Haratifar, 2008
Dissolved oxygen
Protection of flavour after spray-drying
Food (in development)
Inoue and Kato, 2003
Enhancement of probiotic bacteria growth
Functional food (in development)
Bolduc et al., 2006
Food (in development)
Mondal and Lalvani, 2003
Electrode reactions Electrolysis with Coffee membrane © Woodhead Publishing Limited, 2010
Oil
Trans fatty acid
Decrease trans fatty acid content
Water
Electroreductive compounds
Increase purity and quality Food of tap water (in development)
Crandall et al., 2001 Koseki et al., 2003 Mercier, 1999
Increase microbiological quality of water
Food (in development)
Kim et al., 2000 Thompson and Gerson, 1985
Electroreduction of whey protein
Food (in development)
Bazinet et al., 1997a
Production of protein isolate
Food (in development)
Janson and Lewis, 1994
Whey
Whey proteins
Electrodialytic phenomena and associated electromembrane technologies 207
Table 7.1 Electrodialytic phenomena and associated applications in the food, beverage and nutraceutical industries
Table 7.1 Continued Substrate
Membrane reactions Electrodialysis Maple sap with ion-exchange membranes Milk © Woodhead Publishing Limited, 2010
Target compounds Objective
Applications and stage of development
References
Calcium
Food (in development)
Bazinet et al., 2007
Food (in development) Food (industrial scale) Food (in development) Food (industrial scale)
Bolzer, 1985
Casein Minerals
Passion fruit Citric acid juice Sugar beet juice, Minerals sugar cane and molasses Soy tofu whey Minerals
Whey
Magnesium or calcium Lactic acid
To avoid sugar sand formation during maple syrup production Production of acid caseinate Demineralization of milk for further applications Deacidification of fruit juice Demineralization of sugar syrups Use of soy tofu whey as bacteria growth medium Recovery of coagulant agent Production of biological preservative agent
Minerals
Demineralization of whey for further applications
Propionic acid
Production of yeast inhibitor Whey protein separation
Whey
Food (in development) Food (in development) Food and biopharmaceutical (in development) Food (industrial scale) Food (in development) Food (in development)
Hiraoka et al., 1979 Vera Calle et al., 2002, 2003 Chaput, 1979 Ben Ounis et al., 2008 Bazinet et al., 1999b Boyaval et al., 1987 Glassner, 1992 Houldsworth, 1980 Pérez et al., 1994 Boyaval et al., 1993 Amundson et al., 1982 Slack et al., 1986; Stack et al., 1995
208 Separation, extraction and concentration processes
Technology
Electrodialysis with Apple juice bipolar membranes
Potassium
To avoid potassium bitartrate formation
Polyphenol oxidase Inhibition of enzymatic browning
Food (industrial scale)
Audinos et al., 1979, 1985 Escudier et al., 1995 Guérif, 1993
Food (in development)
Quoc et al., 2000, 2006 Tronc et al., 1997, 1998
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Chitosan
Chitosanase
Chitosanase inhibition and Nutraceuticals optimization of chitosane (in development) oligomer production
Lin Teng Shee et al., 2008
Milk
Casein
Production of isolate
Food (in development)
Bazinet et al., 1999a, 2001
Passion fruit juice
Citric acid
Deacidification of fruit juice
Food (in development)
Vera-Calle et al., 2002, 2003
Soybean
Soy proteins
Production of isolate
Food (in development)
Bazinet et al., 1996, 1997b, 1997c, 1998 Mondor et al., 2004 Skorepova and Moresoli, 2007
11S and 7S
Fractionation of 7S and 11S fractions
Food (in development)
Bazinet et al., 2000
Soy proteins
Recovery of soy proteins
Food (in development)
Bazinet et al., 1999b
Calcium or magnesium
Recovery of coagulant agent
Food (in development)
Bazinet et al., 1999b
a-lactalbumin (a-la) Production of a-la and and b-lactoglobulin b-lg enriched fractions (b-lg)
Food (in development)
Bazinet et al., 2004a, 2004b
Soy tofu whey
Whey
Phospholipids
Production of phospholipid Food and nutraceuticals Lin Teng Shee et al., 2005, (in development) 2007 and protein enriched fractions
Electrodialytic phenomena and associated electromembrane technologies 209
Wine
Technology
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Electrodialysis with filtration membranes and electromembrane filtration
Substrate
Target compounds Objective
Applications and stage of development
References
Whey protein
Prodution of isolate
Food (in development)
Bazinet et al., 2004a
Lactic acid
Production of biologic preservative agent
Food and biopharmaceutical (in development)
Norddahl, 1998 Norddahl et al., 2001
Peptide VW
Production of an ACE inhibitor peptidic fraction
Food, nutraceutical and bio-pharmaceutical (in development)
Firdaous et al., 2009 Bazinet and Firdaous, 2009a
Cranberry juice Proanthocyanidins and anthocyanins
Production of antioxidant enriched fruit juices
Food and nutraceutical (in development)
Bazinet et al., 2009 Bazinet and Firdaous, 2009b
Green tea
Catechin
Production of antioxidants Food, nutraceutical and bio-pharmaceutical (in development)
Labbé et al., 2005
Milk
as2-Casein hydrolysate (fraction 183-207)
Bioactive peptide separation
Bargeman et al., 2000, 2002
Snow crab
Protein hydrolysate Production of an anticancer Nutraceutical and peptidic fraction biopharmaceutical (in development)
Doyen et al., 2010
Whey
ACE inhibitor bioactive b-Lactoglobulin peptide separation hydrolysate (fraction 142–148)
Nutraceutical and biopharmaceutical (in development)
Bazinet et al., 2005 Poulin et al., 2006, 2007
Lactoferrin
Food (in development)
Ndiaye et al., 2010
Alfalfa white protein
Separation of lactoferrin from whey
Nutraceutical and biopharmaceutical (in development)
210 Separation, extraction and concentration processes
Table 7.1 Continued
Electrodialytic phenomena and associated electromembrane technologies 211
with filtration membranes (EDFMs), couples size-exclusion capabilities of porous membranes with the charge selectivity of electrodialysis (ED). Because no pressure is applied in the electrodialysis cell, the electric field is the only driving force (Bazinet and Firdaous, 2009b). This technology has already been tested for the separation/purification of various high-added-value bioactive molecules (Table 7.1) such as anticancer peptides and antioxidant polyphenols that will be discussed in more detail. Production of an anticancer peptidic fraction Protein hydrolysates from marine products have interesting bioactive properties such as antithrombotic, antihypertensive (Kim and Mendis, 2006), antioxidant (Amarowicz and Shahidi, 1997), anticancer (Picot et al., 2006) and antimicrobial activities (Beaulieu et al., 2010). Recently, a snow crab by-product hydrolysate, has demonstrated antibacterial inhibition properties against specific strains such as Aeromonas hydrophila, Vibrio vulnificus and Vibrio parahaemolyticus (Beaulieu et al., 2010). Further investigations on the bioactive properties of polypeptides originating from this specific snow crab by-product hydrolysate were performed by Doyen et al. (2010). They carried out an electroseparation of the mixture at pH 9, 6 and 3. An ultrafiltration membrane was placed in an electrodialysis cell on both side of hydrolysate compartment allowing the simultaneous separation and recovery of anionic and cationic peptides. The pH values of the recovery compartments were also maintained at pH 9, 6 and 3 during all the separation process. The authors concluded that a selective separation was obtained by the recovery of two anionic and three cationic peptides in the adjacent compartment. They also tested the peptidic fractions for their anti-cancer activity. The results showed that only one fraction amongst the six fractions produced demonstrated a significant inhibition of growth cancerous lines especially on A549 (lung), BT549 (breast) and PC3 (prostate) cell lines. Furthermore, no anticancer aspect was observed for the feed hydrolysate fractions before separation. Consequently, the authors concluded that the anticancer activity was obtained after purification by electrodialysis with ultrafiltration membranes. Indeed, at a low concentration of 190 mg mL–1, cellular mortality of 75–85% for the A549 cell line, of 80% for the BT549 and of 95–100% for the PC3 cell line were obtained. Although the mechanism of action is still unknown, the future perspectives are very interesting because it is possible to isolate each peptide of the fractions recovered and to test them to determine the peptide(s) sequence(s) of interest. Antioxidant enrichment of cranberry juice Cranberry fruits are a rich source of phenolic phytochemicals including phenolic acids, flavonoids and ellagic acids (Vattem et al., 2005). Many of these compounds have antioxidant activity, equal to or greater than vitamin E (Yan et al., 2002). Cranberry juice can also prevent gastric ulcers caused
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212 Separation, extraction and concentration processes by Helicobacter pylori (Shmuely et al., 2007) and displays potent anticancer activity (Neto et al., 2008). Bazinet et al. (2009) circulated in both compartments on each side of the filtration membrane a volume of cranberry juice (respectively, 450 mL and 200 mL) to concentrate a cranberry juice in its own specific compounds. The total concentrations of proanthocyanidins and anthocyanins increased by 34.8 and 52.9%, respectively in the 200 mL cranberry juice treated by EDFM with a 500 kDa filtration membrane. Moreover, a 18% increase of the antioxidant capacity was measured by means of the oxygen radical absorbance capacity (ORAC) test in the 200 mL of enriched cranberry juice. Moreover, the taste of the enriched cranberry juice was better than the non-treated juice. Based on these results the authors concluded that the production of phenolic antioxidant enriched cranberry juice could be feasible on a large scale and proposed an integrated process flow. According to this process flow, the EDFM process would be directly connected to the bottling process of cranberry juice to produce antioxidant-enriched cranberry juices by a batch process and cranberry juice with very low variation in antioxidant in a continuous process. An advantage of this technology is that compounds of interest would be directly transferred from one juice to another without the need for solvent to extract them first. 7.3.4 Electrodialysis with bipolar membranes A new type of membrane, called a bipolar membrane (BPM), appeared commercially at the end of the 1980s. Bipolar membranes carry out the dissociation of water in the presence of an electric field. These membranes are composed of three parts: an anion-exchange layer, a cation-exchange layer, and a hydrophilic transition layer at their junctions (Mani, 1991). In electrodialysis with bipolar membranes (EDBPMs), cation- and anion-exchange membranes are stacked together with bipolar membranes in an alternating series in an electrodialysis cell and allow its application to numerous products (Table 7.1). Amongst these applications, only the deacidification of fruit juices and the fractionation of 11S-7S soybean protein is developed. Deacidification of passion fruit juices The yellow passion fruit, Passiflora edulis f. flavicarpa, has an intense and special aroma and flavour, which make it a desirable ingredient in the formulation of various food products. However, because of its high acidity, only limited amounts of juice can be added as an ingredient in the formulation of various preparations (Adhikary et al., 1983; Couture and Rousseff, 1992). To overcome this Vera Calle et al. (2002, 2003) tested the deacidification of clarified passion fruit juice by electrodialysis with bipolar membranes (EDBPMs). The stack was equipped with homopolar and bipolar membranes, forming two or three compartments. The reduction of acidity was achieved by increasing the pH from 2.9 to 4.0. This pH limit was chosen to avoid
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Electrodialytic phenomena and associated electromembrane technologies 213
micro-organism growth and spoilage (Vera Calle et al., 2002). In these EDBPM configurations only anions were able to pass through the anionexchange membranes from the juice to the concentrate compartment. The net effect was the extraction of anions, mainly citrate, and their replacement by hydroxyl ions provided by the BPM. Citric acid was formed in the concentrate compartment by citrate ions extracted from juice and protons provided by the second BPM separating the concentrate compartment from the electrode compartments. The anion concentration was similar for the two processes. The inorganic ions were almost eliminated, and 62% of the citrate ions and 48% of the malate ions were removed from the fresh juice. According to the sensory properties of deacidified juices, no significant differences were observed between the deacidified juices and the fresh juice. 11S–7S fractionation The two major reserve soybean proteins, the globulins 7S or b-conglycinin (37–39% of total protein) and 11S or glycinin (31–44% of total protein) have different intrinsic properties leading to different functional properties. From results presented by Bazinet et al. (2000), it appears that during the EDBPM process, the temperature has a large effect on the selective precipitation of the soybean protein fractions, 11S and 7S. Hence, at 27 °C, the precipitation profile of the four protein fractions is situated in a pH range from 6.6 to 4.4, with no possibility of separating any of these fractions. However, at 10 °C, the 11S globulin precipitates at a higher pH than it does at 27 °C, respectively, pH 6.7 compared with 5.9, allowing the fractionation of 11S from the other fractions. Using electroacidification, it is possible to obtain a precipitate solution enriched in the 11S fraction (71.8% of 11S and 10.8% of 7S) and a supernatant solution enriched in the 7S fraction (46.6% of 7S and 4.6% of 11S). EDBPM appears to be a promising technology for fractionation of proteins, by adjusting the electroacidification and solution parameters. This technology does not denature protein (Bazinet et al., 1996), does not use any chemical acids or bases during the process, and produces isolates with a low ash content, which results in a final product of high purity.
7.4 Future trends In this chapter, an overwiew of electromembrane technologies and their associated applications has been presented. Numerous studies have already been performed on various raw matrices (milk, soya, whey, fruit juices) using different cell configurations, but new applications await discovery. Very recently electromembrane processes allowed the separation of bioactive peptides from snow-crab by-products (Doyen et al., 2010) and alfalfa protein hydrolyzates (Firdaous et al., 2009), and the enrichment of cranberry juices in antioxidant compounds (Bazinet et al., 2009). In addition, the low energy © Woodhead Publishing Limited, 2010
214 Separation, extraction and concentration processes cost and constant selectivity during separation of EDFM by absence of transmembrane pressure and, consequently, no membrane fouling, represent an improvement over the conventional separation processes used in the agrifood industries. Furthermore, many acid or alkaline solutions traditionally used in conventional processes are generated in situ by the EDBPM cell configuration. Based on these observations, it appears that electromembrane technologies present several advantages in agri-food industries, in particular, enrichment of nutraceutical products to improve their health benefits, in response to customer demands. In the future, numerous aspects will have to be explored to enhance the understanding of the technologies. Hence, studies on (1) protein/peptide interactions with filtration and ion exchange membranes (IEM) interfaces during treatment/separation of hydrolyzate solutions, and (2) the role of hydrolyzate mineral content on the modification of peptide theoretical charges and influence on their migration rates, would be of major interest. The understanding of both phenomena could lead to the development of new membrane materials and applications, and consequently, to promising ways for new products manufacturing.
7.5 References Adhikary, S.K., Harkare, W.P., Gowindan, K.P. and Nanjundaswamy, A.M. (1983) Deacidification of fruit juices by electrodialysis, Indian J. Technol., 21: 120–123. Amarowicz, R. and Shahidi, F. (1997) Antioxidant activity of peptide fractions of capelin protein hydrolysates, Food Chem., 58: 355–359. Amundson, C.H., Watanawanichakorn, S. and Hill, C.G., Jr. (1982) Production of enriched protein fractions of b-lactoglobulin and a-lactalbumin from cheese whey. J. Food Process Preserv, 6: 55–71. Audinos, R., Roson, J.P. and Jouret, C. (1979) Application de l’électrodialyse à l’élimination de certains composants du jus de raisin et du vin. Connaissance Vigne Vin, 13: 229–239. Audinos, R., Lurton, L. and Moutounet, M. (1985) Advantage of electrodialysis to produce sweetening products from grape. Sci. Aliment., 5: 619–637. Bargeman, G., Dohmen-Speelmans, M., Recio, I., Timmer, M. and Van den Horst, C. (2000) Selective isolation of cationic amino acids and peptides by electro-membrane filtration. Lait, 80: 175–185. Bargeman, G., Houwing, J., Recio, I., Koops, G.H. and Van den Horst, C. (2002) Electromembrane filtration for the selective isolation of bioactive peptides from an as2-casein hydrolysate. Biotechnol Bioeng., 80: 599–609. Bazinet, L., Cossec, C., Gaudreau, H. and Desjardins, Y. (2009) Production of a phenolic antioxidant enriched cranberry juice by an electrochemical process. J. Agric. Food Chem., 57: 10245–10251. Bazinet, L. (2005) Electrodialytic phenomena and their applications in the dairy industry: a review. CRC Critical review in Food Science and Nutrition, 45: 307–326.
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Electrodialytic phenomena and associated electromembrane technologies 215 Bazinet, L., Amiot, J., Poulin, J.-F. and Labbé, D. and Tremblay, A. (2005) Process and system for separation of organic charged compounds. World Patent 082495A1. Bazinet, L. and Firdaous, L. (2009a) Membrane processes and devices for separation of bioactive peptides. Recent Pat. Biotechnol., 3: 61–72. Bazinet, L. and Firdaous, L. (2009b) Applications of electromembrane processes to the production of nutraceuticals or functional foods. In Handbook of membrane research: properties, performance and applications. Gorley, S.V. (Editor). Chemical engineering methods and technology series, Nova Science Publishers Inc., Hauppauge, NY. Chapitre 8, pp. 291–312. Bazinet, L., Ippersiel, D. and Lamarche, F. (1999b) Recovery of magnesium and protein from soy tofu whey by electrodialytic configurations. J. Chem. Technol. Biotechnol., 74(7): 663–668. Bazinet, L., Lamarche, F. and Ippersiel, D. (1998) Comparison of chemical and bipolar membrane electrochemical acidification for precipitation of soybean proteins. J. Agric. Food. Chem., 46: 2013–2019. Bazinet, L., Lamarche, F., Boulet, M. and Amiot, J. (1997a) Combined effect of pH and temperature during electroreduction of whey proteins. J. Agric. Food Chem., 45: 101–107. Bazinet, L., Lamarche, F., Labrecque, R. and Ippersiel, D. (1997b) Effect of KCl and soya protein concentrations on the performance of bipolar membrane electro-acidification. J. Agric. Food. Chem., 45: 2419–2425. Bazinet, L., Lamarche, F., Labrecque, R. and Ippersiel, D. (1997c) Effect of number of bipolar membranes and temperature on the performance of bipolar membrane electroacidification. J. Agric. Food. Chem., 45: 3788–3794. Bazinet, L., Gaudreau, H., Lavigne, D. and Martin, N. (2007) Partial demineralization of maple sap by electrodialysis: impact on syrup sensory and physicochemical quality. J. Sci. Food Agric., 87: 1691–1698. Bazinet, L., Ippersiel, D., Gendron, C., Mahdavi, B., Amiot, J. and Lamarche, F. (2001) Effect of added salt and increase in ionic strength on skim milk electroacidification performances. J. Dairy Res., 68: 237–250. Bazinet, L., Ippersiel, D., Labrecque, R. and Lamarche, F. (2000) Effect of temperature on the separation of soybean 11S and 7S protein fractions during bipolar membrane electroacidification. Biotechnol. Prog., 16: 292–295. Bazinet, L., Ippersiel, D. and Mahdavi, B. (2004a) Effect of conductivity adjustment on the separation of whey protein by bipolar membrane electroacidification. J. Agric. Food Chem., 52: 1980–1984. Bazinet, L., Ippersiel, D. and Mahdavi, B. (2004b) Fractionation of whey protein by bipolar membrane electroacidification. Innov. Food Sci. Emerg. Technol., 5: 17–25. Bazinet, L., Lamarche, F., Ippersiel, D. and Amiot, J. (1999a) Bipolar membrane electro-acidification to produce bovine milk casein isolate. J. Agric. Food Chem., 47: 5291–5296. Bazinet, L., Lamarche, F., Labrecque, R., Toupin, R., Boulet, M. and Ippersiel, D. (1996) Systematic study on the preparation of a food grade soyabean protein. Rep. Can. Electr. Assoc., 9326, U 987, Research and Development; Montréal. Beaulieu, L., Thibodeau, J., Desbiens, M., Saint-Louis, R. Zatylny, C. and Thibaut, S. (2010) Evidence of antimicrobial activities in peptide fractions originating from snow crab (Chionoecetes opilio) by-products. Probiot. Antimicrob. Prot., DOI 10.1007/ s12602-010-9043-6, in press. Ben Ounis, W., Champagne, C. P., Makhlouf, J. and Bazinet, L. (2008) Utilization of tofu whey pre-treated by electromembrane process as a growth medium for Lactobacillus plantarum LB17. Desalination, 229: 192–203. Bolduc, M.P., Bazinet, L., Lessard, J., Chapuzet, J.M. and Vuillemard, J.C. (2006) Electrochemical modification of the redox potential of pasteurized milk and its evolution during storage. J. Agric. Food Chem., 54: 4651–4657.
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216 Separation, extraction and concentration processes Bolzer, R. (1985) Installation and process for the preparation of acid caseinates. United States Patent 4 559 119. Boyaval, P., Corre, C. and Terre, S. (1987) Continuous lactic acid fermentation with concentrated product recovery by ultrafiltration and electrodialysis. Biotech. Lett., 9: 207–212. Boyaval, P., Seta, J. and Gavach, C. (1993) Concentrated propionic acid production by electrodialysis. Enzyme Microb. Technol.; 15: 683–686. Brett, C.M.A. and Oliveira-Brett, A.M. (1994) Fundamentals of kinetics and mechanism of electrode reactions. In Electrochemistry: principles, methods, and applications; Oxford University Press: New York. Chaput, M. (1979) Le lactose, extraction et hydrolyse et déminéralisation du lactosérum. Rev. Lait. Française, 372: 23–26. Couture, R. and Rouseff, R. (1992) Debittering and deacidifying sour orange (Citrus aurantium) juice using neutral and anion exchange resins. J. Food Sci., 57: 380– 384. Crandall, J.J., Mantz, B.W. and Martz, E.C. (2001) Method for generating oxygenated water, United States Patent 6 284 293 B1. Donnan, F.G. (1911) Theory of membrane equilibrium and membrane potential in the presence of non-dialysing electrolytes: A contribution to physical-chemical physiology. Z. Elektrochem. Angewandte Phys. Chem., 17: 572–581. Doyen, A., Beaulieu, L., Saucier, L., Pouliot, Y. and Bazinet, L. (2010) Peptides from a snow crab by-products hydrolysate demonstrated anticancer properties in vitro after simultaneous separation by electrodialysis with ultrafiltration membranes. Submitted to Sep. Pur. Technol. Escudier, J., Saint-Pierre, B., Batlle, J. and Moutounet, M. (1995) Automatically controlled tartric stabilisation of wine by membrane electrodialysis which reduces its conductivity to the desired level. World Patent 9506110-A1. Firdaous, L., Dhulster, P., Amiot, J., Gaudreau, A., Lecouturier, D., Kapel, R., Lutin, F., Vézina, L.-P. and Bazinet, L. (2009) Concentration and selective separation of bioactive peptides from an alfalfa white protein hydrolysate by electrodialysis with ultrafiltration membranes. J. Membr. Sci., 329: 60–67. Gardais, D. (1990) Les procédés électriques de traitement des rejets industriels. In Environnement et Electricité. Electra, Doppee diffusion, Avon, pp. 200–310. Glassner, D. (1992) ED Applications in biotechnology. In Proceedings 10th annual membrane technology planning conference, Business Communications, Norwalk, pp. 158–165. Gonçalves, F., Fernandes, C., Cameira Dos Santos, P. and De Pinho, M.N. (2003) Wine tartaric stabilization by electrodialysis and its assessment by the saturation temperature. J. Food Eng., 59: 229–235. Guérif, G. (1993) Electrodialysis applied to tartaric stabilisation of wines. Rev. Oenol. Techn. Vitic. Œnol., 69: 39–42. Haratifar, S. (2008) The stability of electro-reduced milk lipids. Master of Science Thesis, Faculty of Agricultural and Food Sciences, Université Laval., Québec, Canada, 94p. Hekal, I.M. (1983) Process for the preservation of color and flavour in liquid containing comestibles. United States Patent 4.374.714. Hiraoka, Y., Itoh, K. and Taneya, S. (1979) Demineralization of cheese whey and skimmed milk by electrodialysis with ion exchange membranes. Milchwissenschaft, 34: 397–400. Houldsworth, D.W. (1980) Demineralization of whey by means of ion exchange and electrodialysis. J. Soc. Dairy Technol., 33: 45–51. Inoue, T.T. and Kato, W.K. (2003). Powdery drinks and process for producing the same. International Patent WO03053153. Janson, H.V. and Lewis, M.J. (1994) Electrochemical coagulation of whey protein. J. Soc. Dairy Technol., 47: 87–90. © Woodhead Publishing Limited, 2010
Electrodialytic phenomena and associated electromembrane technologies 217 Janson, H.V., Kiis, A.A., Alekseev, N.G. and Kirm, A.A. (1990) USSR Patent 1 570 694, 1990. Khidirov, S.S. and Merzametov, M.M. (1982) Electrocoagulation of milk proteins, brief communications, Vol.1, Book 2, XXI International Dairy Congress, Moscow, 12–16 July 1982. Kim, C., Hung, Y.C. and Bracket, R.E. (2000) Roles of oxidation–reduction potential in electrolyzed oxidizing and chemically modified water for the inactivation of foodrelated pathogens. J. Food Prot., 63: 19–24. Kim, S.-K. and Mendis, E. (2006) Bioactive compounds from marine processing byproducts. A review. Food Res. Int., 39: 383–393. Koseki, M., Nakagawa, A., Tanaka, Y., Noguchi, H. and Omochi, T. (2003) Sensory evaluation of taste of alkali-ion water and bottled mineral waters. J. Food Sci., 68: 354–358. Labbé, D., Araya-Farias, M., Tremblay, A. and Bazinet, L. (2005) Electromigration feasability of green tea catechins. J. Membr. Sci., 254: 101–109. Lin Teng Shee, F., Angers, P. and Bazinet, L. (2005) Precipitation of cheddar cheese whey lipids by electrochemical acidification. J. Agric. Food Chem., 53: 5635–5639. Lin Teng Shee, F., Angers, P. and Bazinet, L. (2007) Delipidation of a whey protein concentrate by electroacidification with bipolar membranes (BMEA). J. Agric. Food Chem., 55: 3985–3989. Lin Teng Shee, F., Arul, J., Brunet, S. and Bazinet, L. (2008) Performing a three-step process for conversion of chitosan to its oligomers using an unique bipolar membrane electrodialysis system. J. Agric. Food Chem., 56: 10019–10026. Mani, K.N. (1991) Electrodialysis water splitting technology. J. Membr. Sci., 58: 117–138. Mercier, D. (1999) Electrochemical treatment method and device for softening water. United States Patent 5.897.765. Mondal, K. and Lalvani, S.B. (2003) Electrochemical hydrogenation of canola oil using a hydrogen transfer agent. J. Amer. Oil Chem. Soc., 80: 1135–1141. Mondor, M., Ippersiel, D., Lamarche, F. and Boye, J.I. (2004) Production of soy protein concentrates using a combination of electroacidification and ultrafiltration. J. Agric. Food Chem., 52: 6991–6996. Ndiaye, N., Pouliot, Y., Saucier, L., Beaulieu, L. and Bazinet, L. (2010) Electroseparation of bovine lactoferrin from model and whey solutions. Sep. Pur. Technol., 74: 93–99, DOI 10.1016/j.seppur.2010.05.011. Neto, C.C., Amoroso, J.W. and Liberty, A.M. (2008) Anticancer activities of cranberry phytochemicals: an update. Mol. Nutr. Food Res., 52: S18–S27. Norddahl, B. (1998) Fermentative production and isolation of lactic acid. World Patent 9 828 433. Norddahl, B., Eriksen, S. and Pedersen, F.M. (2001) Method for producing lactic acid. World Patent 01 92555-A1. Pérez, A., Andrés, L.J., Alvarez, R., Coca, J. and Hill, C.G. (1994) Electrodialysis of whey permeates and retentates obtained by ultrafiltration. J. Food Proc. Eng., 17: 177–190. Picot, L., Bordenave, S., Didelot, S., Fruitier-Arnaudin, I., Sannier, F., Thorkelsson, G., Bergé, J.P., Guérard, F., Chabeaud, A. and Piot. J.M. (2006) Antiproliferative activity of fish protein hydrolysates on human breast cancer cell lines. Process Biochem., 41: 1217–1222. Poulin, J.-F., Amiot, J. and Bazinet, L. (2006) Simultaneous separation of acid and basic bioactive peptides by electrodialysis with ultrafiltration membrane. J. Biotechnol., 123: 314–328. Poulin, J.-F., Amiot, J. and Bazinet, L. (2007) Improved peptide fractionation by electrodialysis with ultrafiltration membrane: influence of ultrafiltration membrane stacking and electrical field strength. J. Membr. Sci., 299: 83–90.
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218 Separation, extraction and concentration processes Quoc, A., Lamarche, F. and Makhlouf, J. (2000) Acceleration of pH variation in cloudy apple juice using electrodialysis with bipolar membranes. J. Agric. Food Chem., 48: 2160–2166. Quoc, A.L., Mondor, M., Lamarche, F., Ippersiel, D., Bazinet, L. and Makhlouf, J. (2006) Effect of a combination of electrodialysis with bipolar membranes and mild heat treatment on the browning and opalescence stability of cloudy apple juice. Food Res. Int., 39: 755–760. Shaposhnik, V.A. and Kesore, K. (1997) An early history of electrodialysis with permselective membranes. J. Membr. Sci., 136: 35–39. Shmuely, H., Yahav, J., Samra, Z., Chodick, G., Koren, R., Niv, Y. and Ofek I. (2007) Effect of cranberry juice on eradication of Helicobacter pylori in patients treated with antibiotics and a proton pump inhibitor. Mol. Nutr. Food Res., 51: 746–751. Skorepova, J. and Moresoli, C. (2007) Carbohydrate and mineral removal during the production of low-phytate soy protein isolate by combined electroacidification and high shear tangential flow ultrafiltration. J. Agric. Food Chem., 55: 5645–5652. Slack, A.W., Amundson, C.H. and Hill, C.G. (1986) Production of enriched b-lactoglobulin and a-lactalbumin whey protein fractions. J. Food Proc. Pres., 10: 19–30. Stack, F.M., Hennessy, M., Mulvihill, D. and O’Kennedy, B.T. (1995) World Patent 9534216-C1. Swanson, A.M. and Sommer, H.H. (1940) Oxidized flavor in milk: II. The relation of oxidation–reduction potentials to its development. J. Dairy Sci., 23: 597–614. Thompson, B.G. and Gerson, D.F. (1985) Electrochemical control of redox potential in batch cultures of Escherichia coli. Biotechnol. Bioeng., 27: 1512–1515. Tronc, J.-S., Lamarche, F. and Makhlouf, J. (1997) Enzymatic browning inhibition in cloudy apple juice by electrodialysis. J. Food Sci. 62: 75–78. Tronc, J.-S., Lamarche, F. and Makhlouf, J. (1998) Effect of pH variation by electrodialysis on the inhibition of enzymatic browning in cloudy apple juice. J. Agric. Food Chem., 46: 829–833. Vattem, D.A., Ghaedian, R. and Shetty, K. (2005) Enhancing health benefits of berries through phenolic antioxidant enrichment: focus on cranberry. Asia Pac. J. Clin. Nutr., 14: 120–130. Vera Calle, E., Dornier, M., Sandeaux, J., Pourcelly, G. (2002) Deacidification of the clarified passion fruit juice using different electrodialysis configurations. Desalination, 149: 357–361. Vera Calle, E., Ruales, J., Sandeaux, J., Dornier, M., Persin, F., Reynes, M. and Pourcelly, G. (2003) Comparison of different methods for deacidification of clarified passion fruit juice. J. Food Eng., 59: 361–367. Yan, X., Murphy, B.T., Hammond, G.B., Vinson, J.A. and Neto, C.C. (2002) Antioxidant activities and antitumor screening of extracts from cranberry fruit (Vaccinium macrocarpon). J. Agric. Food Chem., 50: 5844–5849.
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Principles of pervaporation for the recovery of aroma compounds 219
8 Principles of pervaporation for the recovery of aroma compounds and applications in the food and beverage industries S. Sahin, Middle East Technical University, Turkey Abstract: The principles and transport mechanism of pervaporation, a membrane separation process in which the components from a liquid mixture permeate selectively through a dense membrane, are described. Information on membrane materials is given and studies on the recovery of aroma compounds by pervaporation, one of its most important applications in the food industry, are reviewed. Sources of further information are listed and future trends explored. Key words: aroma compounds, food, pervaporation, beverages.
8.1 Introduction Processing, especially at high temperatures, may cause considerable physical and/or chemical changes in the aroma compounds of foods. Physical losses may take place during evaporation and chemical changes of aroma compounds may occur owing to heat-induced oxidation or Maillard reactions. For example, in the fruit juice industry, changes in aroma compounds are inevitable since pasteurization of fruit juice is required in order to increase the shelf life of the product. In addition, aroma compounds are highly volatile so that most of them are lost by evaporation during the production of concentrated juice. The change/loss of aroma compounds may affect the final product quality and consumer’s acceptance. Therefore, the aroma compounds must be recovered either from the loss stream or removed before the raw material is subjected to heat treatment (aroma stripping) and added back to the concentrated juice. There are several methods for the extraction that minimize chemical changes and physical losses of aroma compounds. These methods can be
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220 Separation, extraction and concentration processes categorized as: vapor–liquid separations based on distillation/evaporation or partial condensation, separations based on gas injection adsorption and supercritical-fluid extraction. However, these separation techniques have some limitations. Pervaporation, which is a membrane process, is a promising alternative method for aroma recovery because of it has advantages over conventional aroma recovery processes such as high selectivity, low energy consumption, moderate operating temperatures, physical separation mechanism and no additive requirement (Lipnizki et al., 2002a). Compared with other membrane processes like ultrafiltration and reverse osmosis, pervaporation exhibits low fluxes. However, selectivities can be quite high. Therefore, this method is especially suitable for the recovery of highly diluted species. This method can be performed on the juice before the evaporation or pasteurization step thus preventing deterioration or loss of the heat-sensitive aroma compounds during the heat treatment. In addition to aroma recovery, this method can be used for dealcoholization of wine or beer, recovery of high-value-added components or removal of organic pollutants from waste streams, and removal of inhibitors from fermentation broths. It can also be used as an analytical separation technique.
8.2 Principles of pervaporation Pervaporation is a membrane technique in which the components from a liquid mixture are separated by means of partial vaporization through a non-porous permselective membrane. The term ‘pervaporation’ is used in order to emphasize the fact that the permeate undergoes a phase change from liquid to vapor during its transport through the barrier. The pervaporation term was first introduced by Kober (1917). In his study, it was observed that a liquid in a collodion (cellulose nitrate) bag that was suspended in the air, evaporated, although the bag was tightly closed. Kober realized that this phenomenon could be used for separation of liquid mixtures. In pervaporation, a membrane separates an upstream mixture in the liquid state from downstream permeates in the gaseous state. The downstream side is maintained at a reduced pressure to ensure the gaseous state. The liquid feed mixture is circulated in contact with the membrane. The permeate in gaseous state is collected in the liquid state after condensation on a cooled wall. A schematic diagram of a typical experimental set-up for pervaporation is shown in Fig. 8.1. The feed solution is kept in a reservoir inside the temperature-controlled water bath and it is pumped to the upper part of the pervaporation cell. The vacuum is applied to the lower compartment of the cell by use of a vacuum pump in order to provide a pressure gradient across the membrane. The vacuum pressure in the system is controlled by a vacuum controller. The permeate is collected in cold traps by condensation, mostly
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Principles of pervaporation for the recovery of aroma compounds 221 Thermocouple
TC Trap 1
Trap 3
Water bath
PC
Trap 2 Peristaltic Pervaporation pump cell
Vacuum pump
Fig. 8.1 Schematic diagram of a typical pervaporation experimental set-up; TC, temperature control; PC, pressure control (reprinted from Isci et al., 2006, with permission from Elsevier).
with liquid nitrogen at –196 °C. There are usually three traps in the set-up; the first two are connected in parallel in order to conduct the experiments continuously and to maintain the steady-state condition during data collection. The third one is connected in series to the parallel traps as a safety trap in order to ensure that practically no permeate reaches the vacuum pump. The components from the liquid feed mixture permeate selectively through a dense membrane driven by a chemical potential gradient obtained by partial pressure reduction on the permeate side. The affinity between the permeant and the polymer material that constitutes the membrane as well as its mobility through the membrane matrix are important for the transport of the compound. Pervaporation is a complex process involving simultaneous heat and mass transfer. Latent heat is required to pervaporate the liquid. The gradient in partial vapor pressure between the feed and the permeate side of the membrane is maintained by reducing the partial vapor pressure in the permeate side. The heat necessary for the evaporation of the permeate has to be transported through the membrane and this transport of energy is coupled to the transport of mass. The evaporation enthalpy is taken from the sensible heat of the liquid feed mixture, leading to a reduction in feed side temperature.
8.3 Transport mechanism in pervaporation for the recovery of aroma compounds Although a phase change from liquid to vapor takes place, this phase change never occurs directly but occurs via a membrane–solute interaction © Woodhead Publishing Limited, 2010
222 Separation, extraction and concentration processes in pervaporation. Therefore, the separation principle is not based on the vapor–liquid equilibrium. Transport through the nonporous membrane is only by diffusion. The solubility and the diffusivities of the constituents in the membrane are important for separation by pervaporation. Therefore, solute transport within the membrane is not directly affected by the external operating conditions. Ideally, the membrane alone determines the solute transport and the characteristics of the separation and so it is important to create the optimum mass transport conditions to and from the membrane (Schafer et al., 2006). The driving force in pervaporation is the gradient of chemical potential of a solute across the selective membrane (Schafer and Crespo, 2002). Thus, the mass transport in a pervaporation membrane can be described as a function of the chemical potential gradient:
Ji = f (D mi)
[8.1]
where Ji and mi are the mass flux and chemical potential of solute i, respectively. The chemical potential can be stated as a function of the state variables; pressure, composition and temperature:
m = f (P, C, T)
[8.2]
Therefore, the differences in the permeate pressure, feed composition and temperature cause differences in chemical potential. For this reason, there are three methods in pervaporation to establish a suitable chemical potential gradient in the membrane as driving force for the mass transport: (a) Vacuum pervaporation: a hydrostatic pressure difference between the feed and the permeate mixture is established by introducing a vacuum on the permeate side. This method is applicable if the volume of permeating vapor is relatively small or the permeate side pressure is not too low. Otherwise, vacuum pumps of extremely large capacities are required and the pumps consume too much energy (Brüschke, 2006). (b) Thermopervaporation: a chemical potential gradient in the membrane is established by providing a temperature difference between the feed and the permeate side (Strathmann and McDonogh, 1993). (c) Sweeping-gas pervaporation: an inert gas stream on the permeate side of the membrane can be used to remove the component and thus establish a concentration gradient driving force. Vacuum pervaporation is the most commonly used method because a chemical potential gradient is readily obtained by reducing partial pressure on the permeate side. Partial pressures of the permeant components are usually lower than their corresponding saturation pressures. Therefore, they are removed as vapor. The principle of solute transport across the membrane in vacuum pervaporation is shown in Fig. 8.2. A nonporous membrane separates a liquid feed, which is usually close to atmospheric pressure, from the downstream compartment where a vacuum is applied. The solutes (denoted by solid circles)
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Principles of pervaporation for the recovery of aroma compounds 223 Membrane
Liquid
Membrane
Liquid
Vapor
(a)
Vapor
(b) Membrane
Liquid
Vapor
(c)
Fig. 8.2 The principle of solute transport across the membrane in pervaporation. For explanation of panels see text.
are first absorbed by the membrane surface when they are in contact with the membrane owing to the interactions between the solutes and polymer (Fig. 8.2a). When the chemical potential of these components in the feed and in the selective layer of the membrane are equal, a thermodynamic equilibrium is reached. The solutes that are absorbed by the membrane surface create a chemical potential difference across the membrane, which causes a diffusive net flux of solute through the membrane polymer (Fig. 8.2b). Ideally, all the solute particles that have diffused through the membrane are desorbed suddenly and removed by vacuum applied on the downstream side of the membrane. As a result, the solute concentration on the membrane downstream surface remains practically zero, and a maximum concentration gradient between the two membrane surfaces is maintained. If the vacuum is not low enough to desorb all the solutes reaching the membrane downstream surface, the concentration of the solute at the bottom of the membrane is not zero and, consequently, the concentration gradient decreases, as the net flux across the membrane decreases. For the ideal diffusion of solute i across the membrane toward the membrane downstream surface, Fick’s First law applies:
dC J i = Di ÈÍ i ˘˙ Î dz ˚
[8.3]
where Ji is the flux of the permeant i, Ci is the concentration of the permeant i in the polymer membrane, z is the position in the membrane, measured
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224 Separation, extraction and concentration processes from the feed side of the membrane and Di is the diffusion coefficient of the permeant i in the membrane. The diffusion coefficient of a specific component through a polymer can be assumed to be independent of the solute concentration. Then, integration across a membrane, with a thickness Z, yields:
J = Di
(Cimem,f – Cimem,p ) Z
[8.4]
where Cimem,f and Cimem,p are the concentrations of the permeant i inside the membrane at its feed and permeate sides, respectively. If the vacuum applied on the permeate side of the membrane is strong enough, we can assume that all of the solute particles that diffuse through the membrane are removed instantaneously. Then Cimem,p can be taken as zero and equation [8.4] reduces to:
J i = Di
(Cimem,f ) Z
[8.5]
Transport, across the membrane, is generally described by the ‘solutiondiffusion mechanism’ (Strathmann and McDonogh, 1993). Assuming constant solute feed concentration and rapid removal of solute on the membrane downstream, the selective transport of the component ‘i’ from the bulk feed to the downstream compartment can ideally be described by the following steps as: 1. equilibrium partitioning (sorption) of the component i from the liquid phase at the feed solution-membrane upstream surface interphase; 2. diffusion of the absorbed component i through the membrane polymer matrix toward the downstream surface of the membrane; 3. equilibrium partitioning (desorption) of the component i from the membrane downstream surface to the permeate vapor. For the equilibrium partitioning in steps 1 and 3, the sorption (partitioning) coefficient of i in the homogeneous membrane polymer Si can be expressed as: Si =
Cimem,f Cil,bulk
[8.6]
where Cil, bulk is the concentration of component i in the bulk liquid. Combination of equations [8.5] and [8.6] yields the transport model commonly applied in pervaporation:
J i = Si Di
(Cil,bulk ) (C l,bulk ) = Pi i Z Z
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[8.7]
Principles of pervaporation for the recovery of aroma compounds 225
where Pi is the permeability of i in the membrane which is the product of the solubility and diffusivity of the solute i in the membrane. It can be concluded from equation [8.7] that, for an ideal case, the flux of a specific component is directly related to the concentration of the component in the bulk feed. In addition, the permeability of the solute can be calculated from the slope of the flux versus feed concentration graph, if the thickness of the membrane is known. The driving force can also be expressed as a partial pressure gradient of the solute i across the membrane. Then, equation [8.7] becomes:
ÈP f – Pip ˘ DP J i = Si Di ÈÍ i ˘˙ = Pi¢ Í i ˙ ÎZ ˚ Î Z ˚
[8.8]
where Pi¢ is the permeability coefficient (kg m–1 s–1 Pa–1), and Pif and Pip are partial pressures of compound i in the feed and permeate sides, respectively. The widely accepted parameters used to describe the performance of the pervaporation process are permeate flux (Ji), selectivity or separation factor (ai/j) and enrichment factor (bi). The permeate flux Ji is defined as the permeate flow rate per unit of membrane area for a given membrane thickness. The separation factor ai/j is a measure of the selectivity and it indicates the preferential permeation of component i compared with component j. It is calculated from the concentrations of the compounds to be separated, i and j, in the feed and the permeate, respectively:
a i /j =
(Cip /Cil,bulk )
(C pj /C l,bulk ) j
[8.9]
The affinity between the permeant and the polymer material that constitutes the membrane and its mobility through the membrane matrix which are responsible for the transport are important parameters for selectivity. Since different species permeate through the membrane at different rates, a substance at a low concentration in the feed stream can be highly enriched in permeate. The enrichment factor bi represents the capacity of the membrane for concentrating component i and is defined as:
bi =
Cip C l,bulk j
[8.10]
If perfect mixing of the feed and sudden removal of the solutes leaving the membrane downstream face can not be provided, boundary layers develop on both sides of the membrane, affecting the solute transport to and away from the respective membrane surface. If the flux of solute i through the membrane is higher than that through the liquid phase toward the membrane, © Woodhead Publishing Limited, 2010
226 Separation, extraction and concentration processes the solute i is depleted in the liquid phase over the membrane upstream surface, resulting in a liquid solute concentration lower than that in the bulk feed. Because the solute concentration in the liquid at the membrane upstream surface determines the partitioning of the solute into the membrane, the concentration of i in the membrane upstream surface is lower with respect to bulk concentration. Therefore, the concentration gradient across the membrane decreases as does the overall flux. This phenomenon is known as ‘concentration polarization’ and affects the fluxes of compounds of high sorption coefficient, even under turbulent hydrodynamic conditions over the membrane. A similar phenomenon can also be found on the membrane downstream face (Schafer and Crespo, 2002). The concentration polarization strongly depends on the feed flow velocity and on the hydraulic diameter. An increase in the feed flow velocity and a decrease in the hydraulic diameter, decrease the effect of concentration polarization (Borjesson et al., 1996). As long as the diffusive flux is sufficiently high compared with the solute flux across the membrane, the boundary layer does not represent an additional transport resistance and the solute concentration at the membrane surface is equal to that in the bulk (Schafer and Crespo, 2007). However, if the diffusive flux across the liquid boundary layer is lower than the maximum achievable transmembrane flux at the respective feed bulk concentration, the boundary layer detrimentally affects the overall transport of solute from the feed bulk to the permeate side of the pervaporation membrane because the membrane surface concentration of solute i is lower than that of the bulk. Overall mass transfer coefficient kov,i can be expressed as the sum of mass transfer resistances:
Ê 1 ˆ Ê 1 ˆ Ê 1 ˆ Ê 1 ˆ ÁË k ˜¯ = ÁË k ˜¯ + ÁË k ˜¯ + ÁË k ˜¯ ov,i bl,i m,i p,i
[8.11]
where kbl,i is the mass transfer coefficient for the boundary layer, km,i is the mass transfer coefficient for the membrane, and kp,i is the mass transfer coefficient for the permeate side. Although in most instances, it is hard to achieve an ideal case, the mass transfer resistance in the feed boundary layer can be neglected compared with the overall mass transfer resistance if the feed flow velocity is maintained sufficiently high during pervaporation (Karlsson and Tragardh, 1993). Moreover, if a good vacuum is applied on the permeate side of the membrane, everything that diffuses through the membrane is instantaneously pumped away and therefore, mass transfer resistance in the permeate side can also be neglected. Then, the only resistance to mass transfer is within the membrane polymer and equation [8.11] simplifies to:
Ê 1 ˆ Ê 1 ˆ ÁË k ˜¯ = ÁË k ˜¯ ov,i m,i
[8.12]
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Principles of pervaporation for the recovery of aroma compounds 227
Therefore, overall mass flux of component i from the bulk of the feed to the bulk of the permeate can be described by the following equation:
J i = kov,i (Cil,bulk – Cip ) = km,i (Cil,bulk – Cip )
[8.13]
p
Another phenomenon affecting pervaporation performance is temperature polarization. Owing to the phase change from the feed to the permeate side, the membrane acts as a heat sink and a temperature difference exists between the bulk of the feed and the feed side of the membrane. The resulting reduction in the driving force and hence in performance is called ‘temperature polarization’ and this effect is unique for pervaporation processes (Brüschke, 2006). Flux changes linearly with change in concentration but exponentially with change in temperature which means that temperature polarization is more important than concentration polarization. Therefore, the development of high flux pervaporation membranes requires the development of modules in which temperature polarization is effectively reduced. Concentration polarization does not change significantly with changes in composition of the feed. However, the effect of temperature polarization is more important if the concentration of the component to be removed, and thus the flux, is high.
8.4 Selection of membranes for pervaporation in the recovery of aroma compounds A membrane can be simply defined as a permselective barrier made up of polymers between the feed and permeate. Polymers are very large molecules (about 1000 to 100 000 times larger than water molecule) which are manufactured by chemical processes. Polymers gain different characteristics during their design. Generally, there are two types of polymers: glassy polymers and rubbery polymers or elastomers. Their ground state at room temperature determines their distinction (Koops and Smolders, 1991). Glassy polymers have a glass transition temperature above the room temperature and can be divided into three groups: crystalline, semi-crystalline and amorphous polymers. The presence of crystallites has an important affect on all kinds of polymers in terms of tensile strength, elasticity, impact strength, solubility and diffusivity in the polymer. Generally, glassy polymers are hydrophilic (Böddeker and Bengtson, 1991). They preferentially permit the permeation of water owing to the presence of groups in the polymer chain that are able to interact with water molecules. They are used mainly for removal of water from organic solvents and solvent mixtures with an emphasis on azeotropic mixtures. Membranes for the removal of small alcohol molecules like methanol or ethanol are also hydrophilic. Some examples of glassy polymers are cellulose acetate, polysulfone, and polyvinylalcohol.
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228 Separation, extraction and concentration processes Polymers with a glass transition temperature below room temperature are classified as rubbers or elastomers. Elastomers normally present higher chain mobility than glassy polymers and contain rather small, non-polar side groups. Elastomers are flexible and mostly hydrophobic polymers. They preferentially permeate organic substances owing to the lack of strong intermolecular forces, such as hydrogen bonding or dipole–dipole interactions. Since they absorb organic molecules with relatively high fluxes, elastomers are the most suitable polymers for the separation of organic molecules from aqueous solutions (Koops and Smolders, 1991) or from organic–organic mixtures. Some of the examples of hydrophobic polymers are polydimethylsiloxane (PDMS), polyether block polyamide (PEBA), polypropylene (PP), and polyvinylidenefluoride (PVDF) (Koops and Smolders, 1991). Glassy and rubbery membranes differ strongly in terms of flexibility of their polymeric structure. Within the rigid polymeric structure of a glassy membrane, the diffusivity of components is strongly related to their molecular volume, whereas in the flexible region of a rubbery membrane, solute–polymer interactions are much more important than the diffusivities of components. This is why the selectivity of glassy membranes is more diffusion controlled, whereas the selectivity of rubbery membranes is more sorption controlled (Schafer and Crespo, 2002). In pervaporation, nonporous membranes which provide selectivity based on solute–membrane interactions rather than on volatilities are used. The membrane type used in the pervaporation experiments determines the selectivity towards the molecules. Therefore, membrane selection is one of the most important issues in pervaporation. The membrane should have high selectivity and flux for the components to be separated. In addition, it must have chemical, mechanical and thermal stability (Schafer and Crespo, 1997). In pervaporation, the feed side of the membrane is highly swollen in contact with the liquid whereas the permeate side is dry and virtually nonswollen. Thus, a high gradient of swelling exists over the separating layer of the membrane, demanding additional resistance and stability. Pervaporation can be classified as hydrophilic and organophilic pervaporation (Lipnizki et al., 1999). In hydrophilic pervaporation, the target compound water is separated from an aqueous–organic mixture whereas in organophilic pervaporation, the target organic compounds are separated from an aqueous–organic mixture (hydrophobic pervaporation) or from an organic–organic mixture by being preferentially permeated through the membrane (target organophilic pervaporation). Generally, in membrane production for pervaporation processes, composite membrane structure is preferred. Composite membranes are composed of three layers. The first layer is the dense layer responsible for the separation. This layer has to be as thin as possible since the transport through the membrane is by diffusion (0.5–2 mm). The second layer is the porous support layer that has an asymmetric pore structure. It has a thickness of 70–100 mm. Structural polymers with high resistance against chemical attack and good thermal and
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Principles of pervaporation for the recovery of aroma compounds 229
mechanical properties such as polyacrylonitrile, polyetherimide, polysulfone, PVDF form the porous support. The final layer, having a thickness of about 100 mm is a woven or nonwoven textile fabric. Polyester, polyethylene, polyphenylene sulfide and similar fibers are used for the textile carrier layer (Brüschke, 2006). The dense separating layer of hydrophilic membranes is made from various polymers that have high affinity towards water. These polymers contain ions, oxygen functions such as hydroxyl-, ester, or carboxylic moieties or nitrogen as imino- or imido- groups. They must be crosslinked in order to render them insoluble after the coating process. For organophilic membranes, the dense separating layer is formed by crosslinked silicones. The most commonly used organophilic membranes are polydimethylsiloxane (PDMS) or polymethyl octyl siloxane (POMS). The monomer of this silicon rubber is: CH3 Si
CH3
O
[8.14] n
POMS is a modified silicon rubber in which one of the methyl groups of the monomer has been replaced by an octyl group. PDMS or POMS membranes with various porous support layers are commercially available. Zeolite-filled silicon membranes have been developed to increase the selectivity and flux. PEBA formulations have elastic segments consisting of thermoplastic polyamides made flexible by elastomeric polyether links. The modification of PDMS membrane by the introduction of rigid organophilic groups has potential in the pervaporation field for the improvement of separation behavior (Luo et al., 2008). Polyphenylmethylsiloxane–cellulose acetate (PPMS–CA) and polydimethylsiloxane–cellulose acetate (PDMS-CA) membranes are used for concentration of volatile organic compounds such as methanol, ethanol and acetone from aqueous solutions by pervaporation. The hydrophobicity of PPMS membranes is stronger than that of PDMS. In the study of Song and Lee (2005), the tube type alumina substrate was modified with silane coupling agent (perfluoroalkylsilane) to obtain hydrophobic membrane. The surface modified membrane showed much higher flux but lower selectivity than nonporous PDMS membrane for the recovery of esters. However, the selectivity of the membrane was sufficiently high. Olsson et al. (2002) studied the effect of three different POMS membranes, a polymethyloctylsiloxane–polyetherimide (POMS-PEI) and two polymethyloctylsiloxane–polyacrylonitrile (POMS-PAN) membranes having different thicknesses for the separation of alcohols, esters and aldehydes. They concluded that the porous support layer could affect the selectivity considerably and membranes should be designed with a low degree of crosslinking to obtain better separation properties, especially for larger permeants. © Woodhead Publishing Limited, 2010
230 Separation, extraction and concentration processes
8.5 Recovery of aroma compounds by pervaporation and applications in the food and beverage industries One of the most promising applications of pervaporation is the recovery of aroma compounds from aqueous mixtures. Aroma is a complex mixture of hundreds of different volatile organic compounds present at very low concentrations, typically at ppm or ppb levels. Aroma compounds belong to various chemical groups such as alcohols, aldehydes and esters. Recovery of natural flavor and aroma compounds has received much more attention in the food, biotechnology and cosmetic industries because natural aroma compounds are preferred to chemically synthesized ones. In the food industry, flavor concentrates are widely used to compensate for the loss of aroma compounds during processing. Because aroma compounds are highly heat sensitive and loss of aroma compounds during heat treatments is inevitable, they should be separated from food systems and added back to the final product. This can be accomplished either by recovering the lost aromas from the loss stream or by stripping the aromas from the raw material stream before processing. This eventually increases the quality and acceptance of the final product. There have been a large number of studies in this area. Pervaporation has been successfully applied to the recovery of the aroma compounds of wine (Karlsson et al., 1995; Schafer and Crespo, 2007; Schafer et al., 1999) and several fruit juices, such as apple (Bengtsson et al., 1992; Börjesson et al., 1996; Olsson and Tragardh, 1999 and 2001), cashew apple (De Assis et al., 2007), kiwifruit (Cassano et al., 2006), strawberry (Isci et al., 2006), pineapple (Pereira et al., 2002 and 2005; Sampranpiboon et al., 2000), passion fruit (Pereira et al., 2002; 2005), banana (Sampranpiboon et al., 2000), orange (Aroujalian and Raisi, 2007; Shepherd et al., 2002), blueberry (Peng and Liu, 2003), billberries (Diban et al., 2008; Garcia et al., 2008), grape (Rajagopalan and Cheryan, 1995) and pomegranate (Raisi et al., 2008). In addition, pervaporation has been found to be suitable for concentration of tea (Kanani et al., 2003), cocoa (Kattenberg and Willemsen, 2001) and dairy aroma compounds (Baudot and Marin, 1996; Overington et al., 2008) from aqueous solutions and also for dealcoholization of alcoholic beverages (Takacs et al., 2007). It was successfully applied for the recovery of terpenes from waste waters in essential oil industry (Charbit et al., 1997) and for deodorization of waste streams (Souchon et al., 2002). It can also be used as an analytical separation technique for the continuous determination of volatiles (Amador-Hernandez and Castro, 2000). In most of the studies, model feed solutions have been used because of the complexity of conducting the pervaporation experiments with real multicomponent mixtures (Baudot and Marin, 1999; Bengtsson et al., 1992; Börjesson et al., 1996; Garcia et al., 2008; Isci et al., 2006; Kanani et al., 2003; Olsson and Tragardh, 1999; 2001; Peng and Liu, 2003; Pereira et al., 2002 and 2005; Raisi et al., 2008; Rajagopalan and Cheryan, 1995;
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Principles of pervaporation for the recovery of aroma compounds 231
Sampranpiboon et al., 2000). However, there are also some studies in which real food samples are used (Alvarez et al., 2000; Kanani et al., 2003; Pereira et al., 2002 and 2005; Raisi et al., 2008; Rajagopalan and Cheryan, 1995; She and Huang, 2006). Most of the research on aroma recovery by pervaporation has been conducted on a laboratory scale. Studies of scale-up of the pervaporation process for the separation of multi-component mixtures are few. Karlsson et al. (1998) developed a model for the scale-up of pervaporation units for aroma recovery in which the effect of concentration polarization was taken into account with a semi-empirical model. A novel process simulation of pervaporation was developed for multicomponent mixtures (Lipnizki et al., 2002a) and this study was continued with the integration and optimization of hydrophobic pervaporation for the recovery of natural aroma compounds in the food industry (Lipnizki et al., 2002b). Trifunovic et al. (2006) investigated module design aspects of pervaporation using a modified version of an existing pervaporation simulation tool for aroma recovery. By applying the simulation to four aroma compounds: two alcohols (n-butanol, n-hexanol) and two esters (isoamyl acetate and ethyl butyrate), the effect of major process and module design parameters on the performance of a single module has been investigated. In general, for the compounds studied, the permeate composition can be manipulated by changing the module geometry and operating conditions because the effect of these parameters is more significant on esters, which in general have high recovery rates, than for alcohols, which are only marginally affected. 8.5.1 Recovery of fruit juice aromas Apple juice is one of the important fruit juices that pervaporation studies have focused on. Börjesson et al. (1996) investigated the performance of six different pervaporation membranes (PDMS-1060, PDMS-1070 (polydimethyl siloxane + silicalite), PDMS–PT 1100 (polydimethylsiloxane), POMS–PEI (polyoctylmethyl siloxane- polyethermide), POMS–PVDF and PEBA) for the recovery of apple juice aromas. Multicomponent solution containing five esters (ethyl acetate, ethyl butanoate, ethyl- 2-methylbutanoate, isopentyl acetate and hexyl acetate), one aldehyde (trans-2-hexenal) and four alcohols (isobutanol, butanol, isopentanol and hexanol) representing typical aroma compounds of apple juice was used in this study. The best performance was obtained with the PDMS–PT 1100, POMS–PEI and POMS–PVDF membranes. The influence of permeate pressure on the recovery of apple juice aroma compounds by pervaporation had also been examined using the same model solution (Olsson and Tragardh, 2001). A mathematical model, which includes the effect of concentration polarization on the feed side of the membrane and its dependence on permeate pressure, the variation in the separation properties of the membrane with permeate pressure and the properties of the individual permeants, was developed to predict the influence of permeate pressure on separation factor in pervaporation. © Woodhead Publishing Limited, 2010
232 Separation, extraction and concentration processes Moreover, the effect of feed flow velocity on recovery of apple juice aroma compounds by pervaporation was studied (Olsson and Tragardh, 1999). Increasing the feed flow velocity did not increase the recovery of the alcohols significantly because the dominating resistance to mass transfer was in the membrane for the alcohols. However, the recovery of esters and aldehydes was improved by increasing feed flow velocity. The dominating resistance to mass transfer was in the liquid feed boundary layer for the esters studied except ethyl acetate. For the aldehyde and ethyl acetate these two resistances were of the same order. Bengtsson et al. (1992) concentrated the twelve selected flavor compounds in apple juice by pervaporation. It was observed that alcohol had the lowest enrichment factors (2–22), the aldehydes had the medium values (16–67) and the esters had the highest enrichment factors (up to 100). The order of enrichment factors of aroma compounds of various chemical groups was the same in the study of Börjesson et al. (1996). De Assis et al. (2007) found that pervaporation can be used for the concentration of cashew apple juice aroma from cashew pulp. The chromatograms showed an increase in the peak area and also in the number of components in the permeate samples compared with the chromatogram of the cashew juice. Recovery of key flavor components from real flavor concentrates (apple essence, orange aroma and black tea distillate) was achieved by pervaporation in the study of She and Hwang (2006). In this study, both continuous and batch operations were carried out. Generally, the acetates and aldehydes exhibited higher enrichment factors than the alcohols. Significant flavor loss was observed in the pervaporation recovery process. PDMS membrane had higher permeation and loss rates than POMS membrane in batch pervaporation of dilute apple essence. Therefore, it was concluded that POMS membrane with a long operating time should be preferred if highly concentrated products and minimized flavor loss are the major concerns. However, PDMS membrane was found to be better if the time is limited. Rajagopalan and Cheryan (1995) studied pervaporation of model flavor compound of grapes (methyl anthranilate) and of commercial grape essence with several membranes: PDMS–PC (polydimethoxylsiloxane–polycarbonate), PEBA, PDMS-1070). PEBA membrane generally gave higher flux and better selectivity. Flux and selectivity decreased linearly with increase in downstream pressure, but increased with temperature. This relationship was not satisfied when the feed concentration was above 50 ppm where concentration polarization became serious. The presence of ethanol in the feed solution lowered separation factors but increased total flux. Experiments with commercial grape essence confirmed the excellent potential of pervaporation for the production of highly enriched flavors. Pervaporation was found to be a promising technique for the recovery of strawberry aroma compounds (Isci et al., 2006). The experiments were performed using binary aqueous solutions of methyl butyrate (MTB) and
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Principles of pervaporation for the recovery of aroma compounds 233
ethyl butyrate (ETB), ternary mixture of MTB, ETB and water and aqueous model strawberry solution containing six aroma compounds (MTB, ETB, butyl butyrate (BTB), methyl caproate (MTC), ethyl caproate (ETC) and linalool). An increase in both flux and selectivity was observed with an increase in feed temperature or decrease in downstream pressure in binary aqueous solution of MTB (Figs 8.3 and 8.4). Selectivity decreased with increase in feed concentration. Enrichment factors of MTB and ETB in binary and ternary solutions were not very different. However, MTB flux was retarded by half by the presence of ETB in ternary solution. In multicomponent mixtures of aqueous strawberry aroma solution, the selectivities of MTB and ETB were adversely affected by the presence of other compounds, owing to coupling effects. It is shown that flux coupling takes place when a permeant of low diffusivity is dragged through the membrane polymer by a permeant of higher diffusivity resulting in higher fluxes of the slower permeant than expected or the reverse. In the study of Sampranpiboon et al. (2000), pervaporation separation was used to recover aroma compounds from ethyl butanoate and ethyl hexanoate mixtures, which are the most important aroma components of pineapple and banana juice, by using POMS and PDMS membranes. The effect of operating conditions on the separation performance was investigated and in general the POMS membrane was found to be more selective to aroma compounds than PDMS membranes. Moreover, a strong interaction was observed between
Total mass flux (kg m–2 h–1)
0.3
0.2
0.1
0 20
30
40 Temperature (°C)
50
60
Fig. 8.3 Effect of feed temperature on total mass flux at different downstream pressures for 100 ppm methyl butyrate solution: ( ) 4 mbar, ( ) 8 mbar (reprinted from Isci et al., 2006, with permission from Elsevier).
© Woodhead Publishing Limited, 2010
234 Separation, extraction and concentration processes 120
100
Selectivity
80
60
40
20
0 20
30
40 Temperature (°C)
50
60
Fig. 8.4 Effect of feed temperature on selectivity of 100 ppm methyl butyrate solution at different downstream pressures: ( ) 4 mbar, ( ) 8 mbar (reprinted from Isci et al., 2006, with permission from Elsevier).
the two permeating components and permeation of one aroma compound was affected by the presence of the other aroma compound. Pineapple and passion fruit juices were chosen in the study of Pereira et al. (2002) and (2005). Performances of various membrane materials were evaluated with binary and quaternary synthetic aqueous solutions of typical tropical fruit aroma compounds such as ethyl acetate, ethyl butanoate, ethyl hexanoate and 1-octen-3-ol (Pereira et al., 2005). Tests were also carried out using single strength and clarified pineapple juices. Composite flat membranes having selective layers ethylene propylene diene terpolymer (EPDM) or ethylene vinyl acetate copolymer (EVA) and composite hollow fiber membrane having selective layer EPDM were used for comparison. It was concluded that choosing a very selective polymer is advantageous when the organic solute concentration is reduced in the feed. However, more permeable membranes are preferred if the feed concentration is high enough to induce phase separation in the permeate after its condensation. Composite EPDM hollow fiber membrane showed the best performance for pervaporation of synthetic and single-strength pineapple juices. Peng and Liu (2003), evaluated the separation factor of six aroma compounds (1-hexanol, 1-heptanol, trans-2-hexenal, ethyl acetate, linalool and d-limonene) representing typical flavoring ingredients from blueberry juice with a PDMS membrane. The results showed that the separation factor was in the range of 70 to 430, depending on molecule size and polarity of the compounds. No compounds, except 1-heptanol, showed any significant coupling effect in the mixture system. The temperature dependency of component fluxes was © Woodhead Publishing Limited, 2010
Principles of pervaporation for the recovery of aroma compounds 235
expressed by the Arrhenius equation and it was found that the temperature dependency of water flux was the highest whereas that of d-limonene was the lowest. Garcia et al. (2008) studied the concentration of trans-hex-2-en-1-ol, which is one of the major impact aroma compounds of bilberries, from aqueous ethanol using commercial PDMS capillary membranes in pervaporation. The flux of the aroma compound through the membrane was not affected from the variation of the feed flow rate. Therefore, it was concluded that the main mass transfer resistance to trans-hex-2-en-1-ol flux is located in the membrane. The temperature dependency of water, ethanol and trans-hex-2en-1-ol fluxes was expressed by the Arrhenius equation. From the magnitude of calculated activation energies, it was understood that water and ethanol were more influenced by variations in operation temperature than that of the aroma compound. Diban et al. (2008) adapted a previously developed mathematical model to predict the behavior of the separation of a mixture of seven volatile compounds that are characteristic of the aroma of bilberry juice from aqueous solution by pervaporation. It was observed that enrichment factors increased with membrane thickness until an asymptotic value was reached, however, partial fluxes decreased. Shepherd et al. (2002) studied the use of PDMS hollow fibers in orange juice aroma recovery. They used orange juice by product as feed and binary synthetic solutions. The well-spaced module was compared with a hollow fiber not containing spacers and with the transverse flow module with respect to their effect on experimental enrichment factors and mass transfer coefficient values and it was found to be feasible for aroma recovery. Aroujalian and Raisi (2007) also studied the pervaporation of orange juice aroma components. The effects of feed flow rate, feed temperature and permeate pressure on the performance of pervaporation was investigated. Feed flow rate had no significant effect on performance of the process. However, higher flux and enrichment factors were obtained as the feed temperature increased, whereas the increase in permeate pressure resulted in lower flux values. Enrichment factors of ethyl acetate, ethyl butyrate and hexanal increased but that of limonene, linalool and a-terpineol decreased as permeate pressure increased. Aroma compounds of pomegranate juice in multicomponent model solution and real fruit juice have been successfully concentrated by pervaporation (Raisi et al., 2008). The model aroma solution and real pomegranate juice showed similar behavior. The POMS membranes produced higher aroma enrichment factor but lower flux the PDMS membranes. Feed flow rate had no significant effect an either total flux on enrichment factor of aroma compounds, whereas on increase in feed temperature resulted in higher flux and enrichment factor. The effect of permeate pressure on pervaporation performance was related to the properties of the aroma compounds. Alvarez et al. (2000) investigated a combination of membrane processes
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236 Separation, extraction and concentration processes (enzymatic membrane reactor, reverse osmosis and pervaporation) for producing clarified juice and apple juice aroma concentrate. Cassano et al. (2006) performed a study to design an integrated membrane process for the production of concentrated kiwi juice and its aroma recovery. Pervaporation was carried out at different stages in the process. They suggested the use of pervaporation for the removal and enrichment of aroma compounds directly from fresh juice before any concentration process. 8.5.2 Recovery of wine aromas In literature, there are some studies on recovery of wine aromas by pervaporation (Karlsson et al., 1995; Schafer and Crespo, 2007; Schafer et al., 1999). Karlsson et al. (1995) carried out pervaporation with Muscat wine in which eight aroma compounds were identified. In this study, the increase in fluxes with feed temperature was very well described by the Arrhenius equation. The effects of temperature on permeation varied for different aroma compounds. Recovery of aroma compounds from a wine-must fermentation by organophilic pervaporation was studied by Schafer et al. (1999). Pervaporation was performed under fermentation conditions. It was shown that organophilic pervaporation can be highly suitable for the continuous recovery of very complex and delicate aromatic profiles produced during microbial fermentation. Schafer and Crespo (2007) determined the degree of concentration polarization in two different flow-cell configurations during recovery of winemust aroma compounds by pervaporation. It was found that concentration polarization could not be overcome even under turbulent feed flow conditions for compounds having high sorption coefficient for the respective membrane polymer. Dealcoholization of alcoholic beverages including beer and wine is another application of pervaporation. The idea behind this process is the separation of ethanol through hydrophobic membranes much more readily than water. There are many publications in this area (Brüschke, 1990; Escoudier et al., 1988; Kimmerle and Gudernatsch, 1991; Takacs et al., 2007). A problem encountered with this application is that the aroma compounds in the wine or beer are generally much more hydrophobic than ethanol and therefore permeate through the membrane even more readily than ethanol. Vankelecom et al. (1997) used a mixture of eight aromatic compounds in water, one of which was ethanol, in their study. Sorption of aromas in PDMS could be explained by the Hildebrandt solubility parameters. Sorption decreased when zeolites were added to the polymer owing to the cross-linking action of the zeolite. Silicalite-filled PDMS membrane was found to be the best for selective removal of ethanol from the aroma mixture. Therefore, high silicalite loading were recommended for preparation of alcohol-free beverages or removal of ethanol from biofermentations.
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Principles of pervaporation for the recovery of aroma compounds 237
Takacs et al. (2007) studied the reduction of alcohol content of wine by the pervaporation process using organophilic flat composite membrane, PDMS. Alcohol concentrate was obtained in the permeate stream. At higher temperatures, the flux of the permeate was higher and the membrane surface demand was smaller but the separation efficiency was lower. Lower pervaporation temperatures were more favorable to avoid aroma loss. In another study on the reduction of alcohol content of wine by pervaporation, the flux of the permeate increased but the separation efficiency and the separation ability of the membrane decreased as temperature increased (Takacs et al., 2007). In pervaporation of alcoholic beverages, complicated interactions between ethanol and aroma compounds often exist and affect their pervaporation performance. Tan et al. (2005) studied a series of model solutions containing ethanol and six typical aroma compounds in alcoholic beverages. The results showed that the solubility thermodynamics with the feed solution activity coefficient had a dominant effect on the permeability rather than diffusion factor. The presence of aroma compounds decreased the permeability coefficient and the separation factor of ethanol compared with those in ethanol–water binary solutions whereas its effect on fluxes was less. The effect of ethanol feed concentration on mass transfer of each compound was much related to the solubility properties of the compound in ethanol and water. It was observed that little interaction exists between aroma compounds in the membrane during the pervaporation process for dilute solution having a concentration below 1000 ppm. 8.5.3 Recovery of tea, cocoa and coffee aromas The possibility of using pervaporation to recover the tea aroma compounds from tea extract was studied by Kanani et al. (2003). Two different membrane types (POMS and PDMS) were used to recover eight aroma compounds that make a significant contribution to tea aroma. Pervaporation was performed with binary aqueous solutions of aroma compounds, with multicomponent mixture and with an actual tea extract. The actual tea aroma extract showed quite different behavior from the model aroma mixture. The results indicated that pervaporation is an attractive technology for the recovery of tea aroma compounds from tea aroma condensates. However, a wide range of selectivities for different aroma compounds results in alterations in the aroma profile of the permeate. Pilot plant tests for obtaining aroma extracts from cocoa using pervaporation were reported by Kattenberg and Willemsen (2001). The constructed test unit could process up to 500 L of feed, using hollow fiber membranes with a PDMS skin and two-stage condensation for permeate recovery. Application of pervaporation for recovering aroma concentrate from a caffeine- or theobromine-containing food such as coffee or tea and, in particular, cocoa was investigated by Kattenberg et al. (2002). In this study, first an
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238 Separation, extraction and concentration processes aqueous extract was obtained. After incubating the food material at a suitable temperature and for a suitable time, the food extract was pervaporated using a hydrophobic pervaporation membrane to obtain the aroma concentrate. 8.5.4 Recovery of dairy aromas It is also possible to concentrate dairy flavor compounds from aqueous solutions using pervaporation. Baudot and Marin (1996) concentrated two dairy aroma compounds (methylthiobutanoate, diacetyl) in model aqueous solutions by pervaporation through two different membranes (silicalitefilled silicone composite membrane (PDMS 1070) and PEBA homogeneous membrane. Tested membranes showed good selectivity for the extraction of methylthiobutanoate (hydrophobic cheese aroma) at high dilution rate. However, these membranes were less selective for the recovery of diacetyl (2,3-butanedione, hydrophilic butter aroma). For this component, two-stage condensation coupled with pervaporation improved the selectivity of the whole process significantly. Overington et al. (2008) worked with the model feed solution containing nine flavor compounds contributing the flavor of cheese and other dairy products in pervaporation. They studied the effect of membrane type (two types of PDMS and a POMS), feed temperature and permeate pressure. Total flux increased with increase in temperature but decreased with increase in permeate pressure. PDMS membranes gave higher total flux than POMS membrane but the POMS membrane gave higher enrichment factors for the major compounds in the permeate. Esters and ketones passed through the membrane more readily than acids. Diffusion was the controlling mechanism for esters and ketones. Enrichment factors decreased with increasing molecular weight within esters and ketones. However, the effect of molecular weight was more complex for acids and depended on the relative importance of sorption and diffusion mechanisms in the membrane. 8.5.5 Pervaporation as an analytical separation technique Analytical separation techniques play an important role in analytical chemistry. Analytical pervaporation has been successfully used for the continuous determination of volatile analytes or volatile reaction products from complex liquid, semi-solid and solid samples in food and beverage industries (Amador-Hernandez and Castro, 2000). Schafer et al. (2006) investigated the potential of pervaporation combined with an electronic nose based on metal oxide sensors for analyzing wine model solutions. It was shown that target solutes present at minor concentrations can be detected even in the presence of bulk interferents such as ethanol by combining pervaporation with an artificial olfactory system based on metal oxide sensors.
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Principles of pervaporation for the recovery of aroma compounds 239
8.5.6 Other applications of pervaporation for the recovery of aromas Pervaporation can also be used for the extraction of aroma compounds from natural matrixes as an alternative to steam distillation and solvent extraction. Enzymatic pretreatment for degradation of cell membrane is necessary in these processes. Figoli et al. (2006) studied the recovery of bergamot peel oil by pervaporation. There are some studies that combine pervaporation with bioprocessing. Bluemke and Schrader (2001) carried out an integrated bioprocess for production and recovery of the synthesized aroma compounds by interlinking a pervaporation membrane (POMS, PEBA) module with a bioreactor. In this study, they used the fungus Ceratocystis moniliformis which produce ethyl acetate, propyl acetate, isobutyl acetate, isoamyl acetate, citronellol and geraniol. In situ product removal using pervaporation leads to decreased product concentrations in the bioreactor and increased microbial growth rates. Total yield of aroma compounds produced was higher in the integrated bioprocess than with batch cultivation. Böddeker et al. (1997) showed that natural vanillin formed by bioconversion of suitable precursors can be recovered directly from the acidified culture broth by pervaporation at elevated temperature. Large volumes of waste water are obtained in most food processes. Since this liquid is generally released into the environment, it represents both pollution and an economical loss because of the cost of the valuable organic molecules involved. Pervaporation can be used to remove these valuable aroma compounds from waste water. Pierre et al. (2001) carried out non-dispersive solvent extraction of three sulfur aroma compounds, dimethyldisulfide, dimethyltrisulfide and S-methyl thiobutanoate, from very dilute aqueous solutions representing real effluent. Mass transfer fluxes obtained experimentally by membrane-based solvent extraction were greater for the three aroma compounds than those obtained by pervaporation. Souchon et al. (2002) used pervaporation process for deodorization of cauliflower blanching effluent in order to reduce its volatile organic compounds content and to recover the valuable food flavoring fraction. Pervaporation has been performed on three sulfur compounds of the cauliflower odor. Pervaporation was successfully applied for the recovery of linalool and eucalyptol which are the most frequent terpenes in wastewaters from the essential oil industry (Charbit et al., 1997).
8.6 Sources of further information and future trends Pervaporation theory and principles, separation characteristics, thermodynamics and polymer materials for membranes are covered in the book edited by Huang (1991), in which the chapter by Böddeker and Bengtson (1991) deals with the separation of organics from water by pervaporation, and that
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240 Separation, extraction and concentration processes by Koops and Smolders (1991) gives information about membranes used in pervaporation. Strathmann and McDonogh (1993) studied mass transfer in pervaporation, characterization of pervaporation membranes and use of pervaporation in biotechnology. Brüschke (2006) studied the transport resistances and principles of pervaporation, membranes and the application of pervaporation in the chemical industry. Mass transport phenomena during the recovery of volatile compounds by pervaporation were investigated by Schafer and Crespo (2002). Pereira et al. (2006) reviewed the applications of pervaporation for the recovery of volatile aroma compounds from fruit juices. The studies on scale-up of the pervaporation process are limited and mostly based on the separation of model solutions. Therefore, pervaporation studies using real food systems should be increased in future. In addition, further study is required on scale-up. Studying pervaporation in large-scale systems provides insight into the industrial application of this method. Future trends are towards the combination of pervaporation with other processes in the food process line.
8.7 References Alvarez S, Riera F A, Alvarez R, Coca J, Cuperus F P, Bouwer S T, Boswinkel C, van Gemert RW, Veldsink JW, Giorno L, Donato L, Todisco S, Drioli E, Olsson J, Tragardh G, Gaeta S N and Panyor L (2000), ‘A new integrated membrane process for producing clarified apple juice and apple juice aroma concentrate’, J Food Eng, 46, 109–125. Amador-Hernandez J and de Castro MDL (2000), ‘Pervaporation: a useful tool in food analysis’, Food Chem, 68, 387–394. Aroujalian A and Raisi A (2007), ‘Recovery of volatile aroma components from orange juice by pervaporation’, J Membr Sci, 303, 154–161. Baudot A and Marin M (1996), ‘Dairy aroma compounds recovery by pervaporation’, J Membr Sci, 120, 207–220. Bengtsson E, Tragardh G and Hallstrom B (1992), ‘Concentration of apple juice aroma from evaporator condensate using pervaporation’, Lebensm Wiss Technol, 25, 29–34. Bluemke W and Schrader J (2001), ‘Integrated bioprocess for enhanced production of natural flavors and fragrances by Ceratocystis moniliformis’, Biomol Eng, 17, 137–142. Borjesson J, Karlsson H O E and Tragardh G (1996), ‘Pervaporation of a model apple juice aroma solution: comparison of membrane performance’, J Membr Sci, 119, 229–239. Böddeker K and Bengtson G (1991), ‘Selective pervaporation of organic from water’, in R. Y. M. Huang, Pervaporation membrane separation processes, Elsevier, Amsterdam, The Netherlands, 437–460. Böddeker K W, Gatfield I L, Jahnig J, Schorm C (1997), ‘Pervaporation at the vapor pressure limit: Vanillin’, J Membr Sci, 137, 155–158. Brüschke H E A (1990), ‘Removal of ethanol from aqueous streams by pervaporation’, Desalination, 77, 323–329. Brüschke H E A (2006), ‘State-of-art of pervaporation process in the chemical industry’, in S. P. Nunes and K. V. Peinemann, membrane technology in the chemical industry, Wiley-VCH, USA, 151–202.
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Principles of pervaporation for the recovery of aroma compounds 241 Cassano A, Figoli A, Tagarelli A, Sindona G, Drioli E (2006), ‘Integrated membrane process for the production of highly nutritional kiwifruit juice’, Desalination, 189, 21–30. Charbit G, Charbit F and Molina C (1997), ‘Study of mass transfer limitations in the deterpenation of waste waters by pervaporation’, J Chem Eng Jpn, 30, 382–387. De Assis A V R, Bizzo H R, da Matta V M, Cabral L M C (2007), ‘Recovery of aroma compounds of cashew apple fruit (Anacardium accidentale L.) by pervaporation’, Ciencia E Technologia De Alimentos, 27, 349–354. Diban N, Urtiaga A and Ortiz I (2008), ‘Recovery of key components of bilberry aroma using a commercial pervaporation membrane’, Desalination, 224, 34–39. Escoudier J L, Le Bouar M, Moutounet M, Jouret C and Barillere J M (1988), ‘Application and evaluation of pervaporation for production of low alcohol wines’, in R. Bakish, Proceedings of the Third International Conference on Pervaporation Processes in the Chemical Industry, Bakish Materials Corporation, Englewood Cliffs, NJ, USA, 387–397. Figoli A, Tagarelli A, Mecchia A, Trotta A, Cavaliere B, Lavecchia R, Sindona G and Drioli E (2006), ‘Enzyme-assisted pervaporative recovery of concentrated bergamot peel oils’, Desalination, 199, 111–112. Garcia V, Diban N, Gorri D, Keiski R, Urtiaga A and Ortiz I (2008), ‘Separation and concentration of bilberry impact aroma compound from dilute model solution by pervaporation’, J Chem Technol Biotechnol, 83, 973–982. Huang R Y M (1991), Pervaporation membrane separation processes, Elsevier Science Publishers, Amsterdam, The Netherlands. Isci A, Sahin S and Sumnu G (2006), ‘Recovery of strawberry aroma compounds by pervaporation’, J Food Eng, 75, 36–42. Kanani D M, Nikhade B P, Balakrishnan P, Singh G and Pangarkar V G (2003), ‘Recovery of Valuable Tea Aroma Components by Pervaporation’, Ind Eng Chem Res, 42, 6924–6932. Karlsson H O E and Tragardh G (1993), ‘Aroma compound recovery with pervaporationfeed flow effects’, J Membr Sci, 81, 163–171. Karlsson H O E, Loureiro S and Tragardh G (1995), ‘Aroma compound recovery with pervaporation-Temperature effects during pervaporation of a muscat wine’, J Food Eng, 26, 177–191. Karlsson H O E, Tragardh G and Olsson J (1998), ‘The performance of pervaporative aroma recovery units: Process simulations’, Sep Sci Technol, 33, 1629–1652. Kattenberg H R and Willemsen J H A (2001), ‘Aroma extracts from cocoa’, Manuf Confect, 82, 73. Kattenberg H R, Willemsen J H A, Starmans D A J, Hoving H D and Winters MGM (2002), ‘Method for recovering aroma concentrate from caffeine- or theobrominecomprising food base material’, US Patent Office Journal, No. 1139792. Kimmerle K and Gudernatsch W (1991), ‘Pilot dealcoholization of beer by pervaporation’, in R. Bakish, Proceedings of the Fifth International Conference on Pervaporation Processes in the Chemical Industry, Bakish Materials Corporation, Englewood Cliffs, NJ, USA, pp. 291–307. Kober P A (1917), ‘Pervaporation, perstillation, and percrystallisation’, J Am Chem Soc, 39, 944–948. Koops G H and Smolders C A (1991), ‘Estimation and evaluation of polymeric materials for pervaporation membranes’, in R. Y. M. Huang, Pervaporation membrane separation processes, Elsevier, Amsterdam, The Netherlands, 253–278. Lipnizki F, Hausmanns S, Ten P, Field R W and Laufenberg G (1999), ‘Organophilic pervaporation: prospects and performance’, Chem Eng J, 73, 113–129. Lipnizki F, Olsson J and Tragardh G (2002a), ‘Scale-up of pervaporation for the recovery of natural aroma compounds in the food industry. Part 1: Simulation and performance’, J Food Eng, 54, 183–195.
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242 Separation, extraction and concentration processes Lipnizki F, Olsson J and Tragardh G (2002b), ‘Scale-up of pervaporation for the recovery of natural aroma compounds in the food industry. Part 2: Integration and optimization’, J Food Eng, 54, 197–205. Luo Y, Tan S J, Wang H, Wu F W, Liu X M, Li L, Zhang Z B (2008), ‘PPMS composite membranes for the concentration of organics from aqueous solutions by pervaporation’, Chem Eng J, 137, 496–502. Olsson J and Tragardh G (1999), ‘Influence of feed flow velocity on pervaporative aroma recovery from a model solution of apple juice aroma compounds’, J Food Eng, 39, 107–115. Olsson J and Tragardh G (2001), ‘Pervaporation of volatile organic compounds from water. I. Influence of permeate pressure on selectivity’, J Membr Sci, 187, 23–37. Olsson J, Tragardh G and Lipnizki F (2002), ‘The influence of permeant and membrane properties on mass transfer in pervaporation of volatile organic compounds from dilute organic solutions’, Sep Sci Technol, 37, 1199–1223. Overington A, Wong M, Harrison J and Ferreira L (2008), ‘Concentration of dairy flavour compounds using pervaporation’, Int Dairy J, 18, 835–848. Peng M and Liu S X (2003), ‘Recovery of aroma compounds from dilute model blueberry solution by pervaporation’, J Food Sci, 68, 2706–2710. Pereira C C, Ribeiro C P, Nobrega R, Borges C P (2006), ‘Pervaporative recovery of volatile aroma compounds from fruit juices’, J Membr Sci, 274, 1–23. Pereira C C, Rufino J M, Habert A C, Nobrega R, Cabral L M C and Borges C P (2002), ‘Membrane for processing tropical fruit juice’, Desalination, 148, 57–60. Pereira C C, Rufino J R M, Habert A C, Nobrega R, Cabral L M C and Borges C P (2005), ‘Aroma compounds recovery of tropical fruit juice by pervaporation: membrane material selection and process evaluation’, J Food Eng, 66, 77–87. Pierre F X, Souchon I, Marin M (2001), ‘Recovery of sulfur aroma compounds using membrane-based solvent extraction, J Membr Sci, 187, 239–253. Raisi A, Aroujalian A, Kaghazchi T (2008), ‘Multicomponent pervaporation process for volatile aroma compounds recovery from pomegranate juice’, J Membr Sci, 322, 339–348. Rajagopalan N and Cheryan M (1995), ‘Pervaporation of grape juice aroma’, J Membr Sci, 104, 243–250. Sampranpiboon P, Jiraratananon R, Uttapap D, Feng X and Huang R Y M (2000), ‘Separation of aroma compounds from aqueous solutions by pervaporation using polyoctylmethyl siloxane (POMS) and polydimethyl siloxane (PDMS) membranes’, J Membr Sci, 174, 55–65. Schafer T and Crespo J G (1997), ‘Recovery of aromas by pervaporation’, ESMST XIVth Annual Summer School, Membrane integration in clean processes. Lisbon Instituto: Superior Tecnico, IST. Schafer T and Crespo J G (2002), ‘Mass transport phenomena during the recovery of volatile compounds by pervaporation’, in J. Welti-Chanes, J. F. Velez-Ruiz and G.V. Barbosa-Canovas, Transport phenomena in food processing, CRC Press, Florida, USA, 247–263. Schafer T and Crespo J G (2007), ‘Study and optimization of the hydrodynamic upstream conditions during recovery of a complex aroma profile by pervaporation’, J Membr Sci, 301, 46–56. Schafer T, Bentson G, Pingel H, Boddeker K W and Crespo J P S G (1999), ‘Recovery of aroma compounds from a wine-must fermentation by ogranophilic pervaporation’, Biotechnol Bioeng, 62, 412–421. Schafer T, Serrano-Santos M B, Rocchi S and Fuoco R (2006), ‘Pervaporation membrane separation process for enhancing the selectivity of an artificial olfactory system (“electronic nose”)’, Anal Bioanal Chem, 384, 860–866.
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Principles of pervaporation for the recovery of aroma compounds 243 She M and Hwang S-T (2006), ‘Recovery of key components from real flavor concentrates by pervaporation’, J Membr Sci, 279, 86–93. Shepherd A, Habert A C and Borges C P (2002) ‘Hollow fiber modules for orange juice aroma recovery using pervaporation’, Desalination, 148, 111–114. Song K H and Lee K R (2005), ‘Pervaporation of flavors with hydrophobic membrane’, Korean J Chem Eng, 22, 735–739. Souchon I, Pierre F X, Athes-Dutour V and Marin M (2002), ‘Pervaporation as a deodorization process applied to food industry effluents: recovery and valorisation of aroma compounds from cauliflower blanching water’, Desalination, 148, 79–85. Strathmann H and McDonogh R M (1993), ‘Use of pervaporation in biotechnology’, in J. A. Howell, V. Sanchez and R. W. Field, Membranes in bioprocessing: theory and applications, Chapman and Hall, Cambridge, UK, 293–329. Takacs L, Vatai G, Korany K (2007), ‘Production of alcohol free wine by pervaporation’, J Food Eng, 78, 118–125. Tan S J, Li L, Xiao Z Y, Wu Y T and Zhang Z B (2005), ‘Pervaporation of alcoholic beverages – the coupling effects between ethanol and aroma compounds’, J Membr Sci, 264, 129–136. Trifunovic O, Lipnizki F and Tragardh G (2006), ‘The influence of process parameters on aroma recovery by hydrophobic pervaporation’, Desalination, 189, 1–12. Vankelecom I F J, Beukelaer S D and Uytterhoeven J B (1997), ‘Sorption and pervaporation of aroma compounds using zeolite-filled PDMS membranes’, J Phys Chem B, 101, 5186–5190.
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244 Separation, extraction and concentration processes
9 Advances in membrane-based concentration in the food and beverage industries: direct osmosis and membrane contactors E. Drioli and A. Cassano, Institute on Membrane Technology, ITM-CNR, Italy Abstract: The process fundamentals and the role of operating conditions involved in direct osmosis (DO) and membrane contactors, such as osmotic distillation (OD) and membrane distillation (MD) are reviewed. Their main applications in concentration of liquid food and limitations compared with those of conventional de-watering processes – such as thermal evaporation, cryoconcentration and reverse osmosis – are discussed. The advantage lies in allowing very high concentrations (up to 65 °Brix) to be reached under atmospheric pressure and low temperature, thus avoiding thermal and mechanical damage. Key words: membrane contactors, direct osmosis, osmotic distillation, membrane distillation, food industry.
9.1 Introduction Epidemiological studies report that high consumption of fruit and vegetables is associated with a reduced risk of free radical related oxidative damage and diseases, such as various types of cancer, and cardiovascular or neurological diseases (Collins and Harrington, 2002). The world market of these products has remarkably increased in recent years and consumers have addressed their interest towards products with a natural fresh taste, that are high quality and additive-free. In order to meet these demands, the food industry has focused on the development of processed items with increased shelf-life able to retain as much as possible the essence of the fresh fruit, as well as colour, aroma, nutritional value and structural characteristics.
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Advances in membrane-based concentration 245 In the traditional methods of processing liquid foods, the concentration operation is particularly important for the production of intermediates that can be easily handled, packaged, stored and shipped before their reutilization for a final transformation (Luh et al., 1986) and also as a step which precedes total dehydration. Moreover, the removal of water reduces the growth of micro-organisms, thus increasing the shelf-life of the liquid foods; fruit juice concentrates, because of their low water activity, have a higher stability than single-strength juices. This chapter briefly describes the existing methods for the concentration of liquid foods (such as thermal evaporation, cryoconcentration and membrane concentration) and then introduces novel membrane processes, including direct osmosis (DO), membrane distillation (MD) and osmotic distillation (OD), which offer improved performances in terms of product quality, reduced energy consumption and environmental impact, and high effectiveness.
9.2 Conventional technologies in the food and beverage industries 9.2.1 Thermal evaporation In thermal evaporation, water is removed from liquid foods as vapour. Falling film evaporators and centrifugal evaporators are most commonly used for this purpose. It is known that the heat required to perform the evaporation results in some ‘cooked’ notes recognized as off-flavours, loss of most volatile aroma compounds, and colour degradation with a consequent remarkable qualitative decline (Maccarone et al., 1996). Partial recovery of aroma compounds may be achieved by cold-water condensation and rectification in a still located on the evaporator, resulting in concentrates that are acceptable to consumers, but still far from fresh products. An additional drawback in the use of thermal evaporation is the high energy demand, despite the use of energy-saving systems. The production of superior quality concentrates, especially fruit juices, is a very important goal and considerable R&D efforts have been devoted to the development of nonthermal concentration techniques. These methods include: freeze concentration systems (cryoconcentration) and membrane processes. 9.2.2 Cryoconcentration Freeze concentration has long been considered a feasible prospect for the selective removal of water from liquid foods. In this instance, water is removed as ice rather than as vapour. The product is cooled below its freezing point allowing water to form crystals which are then removed from the concentrate by centrifugation. This method guarantees an excellent organoleptic quality: the operating temperatures are sufficiently low to avoid chemical and
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246 Separation, extraction and concentration processes biochemical reactions and volatile compounds are completely retained. Consequently, the concentrated product is characterized by higher quality than that processed by thermal evaporation (Aider and de Halleux, 2008). However, there are some problems in this process which make it unpractical. The main drawback is that the achievable concentration (40–45 °Brix) is lower than the values obtained by evaporation (60–65 °Brix). Furthermore, fine ice crystals produced by rapid freezing cannot be separated from the residual liquid and the process is not suitable for the treatment of highly pulped liquids. Remarkable energy consumption, high equipment investments and operational costs are additional drawbacks (Jariel et al., 1996). 9.2.3 Pressure-driven membrane processes Pressure-driven membrane processes such as microfiltration (MF), ultrafiltration (UF), nanofiltration (NF) and reverse osmosis (RO) are well established systems in the food industry for concentration, fractionation and purification of liquid foods and for wastewater treatment (Table 9.1). These processes offer significant advantages over traditional food processing technologies in terms of low operation temperatures, no special chemicals required, possibility of automation, and no phase-change involved. In addition, they are simple in concept and operation and characterized by low energy consumption. MF membranes have a symmetric microporous structure with a pore size in the range 0.1–10 mm. They are essentially used for the separation of fine particles, micro-organisms and emulsion droplets from fluids. Particles are rejected mainly by means of sieving mechanisms although the separation is affected also by interactions between the membrane surface and the treated solution. The hydrostatic pressure difference used as a driving force is in the range from 0.5 to 4 bar. Cell harvesting, clarification of fruit juices, must and wine, wastewater treatment, bacteria and particulate turbidity reduction, separation of whey proteins and separation of oil–water emulsions are typical consolidated applications (Cheryan, 1998). UF membranes have a pore size of 1–100 nm and are capable of retaining species of molecular weight 300–500 000 daltons. Typical rejected species include lipids, proteins, polysaccharides, colloids and suspended solids whereas solutes with small molecules such as vitamins, salts and sugars, flow through the membrane together with water. Operating pressures are 2–10 bar (Mulder, 1998). In the food and dairy industry, UF is commonly applied to clarify, concentrate, decolorize and fractionate juices, sugar solutions, proteins and dairy and grain milling products. The membranes used in NF are characterized by a charged surface and equivalent pore diameters in the range 1–3 nm. They are mainly used for the separation of multivalent ions and uncharged organic molecules having molecular weights greater than 1 kDa. Operating
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Advances in membrane-based concentration 247 Table 9.1 Commercial application of pressure-driven membrane processes in food processing Unit operation Industrial sector
Application
Microfiltration Fruit and vegetables
Fruit and vegetable juice clarification
Beverages Dairy
Ultrafiltration
Fruit and vegetables Beverages Vegetables Dairy
Fats and oils Nanofiltration
Fruit and vegetables Dairy Beverages
Reverse osmosis
Cold sterilization of beer Clarification of wine, beer and vinegar Bacterial purification of milk and whey Whey defatting and high-quality protein recovery Clarification of fruit juice Clarification of wine Recovery of potato starch and proteins Fractionation/concentration of whey in manufacture of whey protein concentrate Fractionation/concentration of milk in manufacture of processed and natural cheese Recovery of oilseed protein and oil from processing water Fruit and vegetable juice concentration Concentration and demineralization of milk Demineralization of whey Wastewater treatment
Fruit and vegetables
Fruit and vegetable juice concentration Sugar concentration and recovery from rinse water, treatment of water for reuse Recovery of flavours, fragrances, pectins and proteins Beverages Recovery of alcohol from wine Sugar and sweeteners Preconcentration of maple syrup Sugar concentration and recovery from rinse water, treatment of water for reuse Dairy Concentration of milk and whey All possible Treatment of boiler concentrate prior to recycle to boiler
pressures in NF are lower than those in RO therefore separation occurs at low energy consumption (21% less than RO). The concentration of must, the recovery of aromas from fruit juices, the demineralization of whey and the treatment of wastewater from beverage production are typical applications of NF in the food and dairy industry (Warczok et al., 2004). In RO, low molecular weight compounds, such as salts and sugars are separated from a solvent, usually water. The particle size range for application of RO is approximately 0.1–1 nm and a complete separation is achieved with solutes of molecular weight greater than 300 Da. Typical operating pressures are 10–100 bar. Food processing applications of RO include: concentration
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248 Separation, extraction and concentration processes of fruit and vegetable juices, concentration of milk and whey for cheese production, pretreatment of boiler water, recovery of sugars and acids from rinse waters, recovery of sweet potato stillage, recovery of flavours, pectins and proteins, removal of alcohol from wine, and water softening (Ho and Sirkar, 1992). The advantages of RO over traditional evaporation are the lower thermal damage to the product, increase in aroma retention, simple system design, lower energy consumption and lower equipment costs. However, RO cannot reach retentate concentrations higher than 25–30 °Brix with a single-stage system owing to the limitation of high osmotic pressure. Furthermore, fouling and concentration polarization are still important drawbacks of RO and other pressure-driven membrane processes in liquid food processing (Jiao et al., 2004).
9.3 Direct osmosis and applications in the food and beverage industries 9.3.1 Process fundamentals In DO, an osmotic pressure gradient across a semi-permeable membrane is established by using an osmotic agent solution in order to remove water from a feed solution. The osmotic agent should be characterized by high solubility in water (low water activity), non-toxicity, high superficial tension, inertia towards the flavour, odour and colour of the foodstuff, low volatility and viscosity, and non-permeability across the membrane. The most frequently employed constituents as osmotic agents are: sodium chloride, calcium chloride, glucose, sucrose, glycerol, cane molasses or corn syrup. Unlike RO the pressure difference across the membrane in DO is negligible and the flux depends on the difference in water activities. The hydraulic pressures required to pump the feed and the osmotic solution over the membrane surface are of about 2 bar (Milleville, 1990). Consequently, fouling phenomena are negligible and only concentration polarization occurs. Additional advantages are low energy consumption, simplicity, modularity, constant permeate flux in time, possibility to treat solutions with a high level of suspended and dissolved solids, and high achievable concentrations. The membranes used in DO are semipermeable membranes similar to those used in RO and NF processes. They are highly selective; however, the diffusion of a small amount of stripping solution cannot be completely avoided. 9.3.2 Effect of operating conditions on the direct osmosis (DO) flux Petrotos et al. (1998, 1999) studied the effect of several membranes and operating parameters on the performance of DO in the concentration of
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Advances in membrane-based concentration 249 tomato juice. NaCl, CaCl2, Ca(NO3)2, glucose, sucrose and polyethylene glycol (PEG 400) were used as stripping solutions. The initial concentration of tomato juice was in the range 4.3–11.7 °Brix. These studies suggested that salts, and especially NaCl, produced improved fluxes compared with carbohydrates and PEG 400, owing to their lower viscosity. An increasing of the stripping solution concentration also positively affected the flux (through an increasing of the driving force). On the contrary, an increasing of the feed concentration resulted in lower water flux owing to viscosity and osmotic pressure increase. The lower viscosity of the osmotic agents resulted in reduced polarization phenomena on the stripping solution side and increased fluxes. Increasing the temperature from 26 to 60 °C enhanced permeate flux by 64%, this phenomenon can be attributed to the reduction in the viscosity and to the increase in the diffusion coefficient of both feed and stripping solution. An increase in feed flow rate did not improve the permeate flux during the DO process. The pretreatment of the juice by MF or UF was also found to have a significant effect on the osmotic fluxes (Petrotos et al., 1998). In particular, the ultrafiltered juice gave a 39% increase in flux compared with the untreated juice. Finally, the DO flux is strongly affected by membrane thickness. In particular, a reduction of the membrane support layer determines an increase of the permeate flux (Beaudry and Lampi, 1990). 9.3.3 DO applications Osmotek Inc. (Corvallis, OR, USA) developed the first DO industrial approach as a pretreatment step for RO wastewater treatment. Several studies concerning the fruit juice concentration by DO are reported (Beaudry and Lampi, 1990; Herron et al., 1994; Milleville, 1990; Petrotos and Lazarides, 2001; Petrotos et al., 1998, 1999; Wrolstad et al., 1993). Beaudry and Lampi (1990) reported the performance of the Osmotek DO process in the concentration of fruit juices by using thin-film composite RO membranes with an overall thickness of 25–85 mm and a nominal molecular weight cut off (NMWCO) of 100 Da. These membranes rejected more than 99.9% of the total acidity and colour in orange and red raspberry juice. Herron et al. (1994) developed a simplified DO apparatus for several kinds of liquid food samples including orange and raspberry juice. Maximum osmotic fluxes of 5–6 L m–2 h–1 were obtained through 100-Da membranes by using fructose/glucose solutions as osmotic agent. The concentrated juice showed a better quality than that produced by a conventional vacuum evaporator. Stripping solutions diluted during juice concentration can be reconcentrated by using thermal evaporation: therefore, only the osmotic agent is exposed to high temperatures preserving the quality of concentrated juice. RO was also proposed as a method for the recovery of stripping solutions (Karode et al., 2000). Petrotos and Lazarides (2001) developed a new apparatus for the concentration of fruit juices with an osmotic cell of special configuration
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250 Separation, extraction and concentration processes to promote turbulence. The apparatus was equipped with RO flat-sheet membranes to concentrate tomato juice up to 16 °Brix at room temperature. Average fluxes of 4.5 kg m–2h–1 were obtained.
9.4 Membrane contactors and applications in the food and beverage industries Membrane contactors are devices in which a microporous membrane acts as a barrier between two phases permitting gas/liquid or liquid/liquid mass transfer of the components without dispersion of one phase into the other. Fluids to be contacted flow on the opposite side of the membrane and a fluid/fluid interface is located at the mouth of each membrane pore. Mass transfer occurs by diffusion across the interface as in conventional contacting equipment. Unlike conventional membrane operations, the driving force for separation is a concentration or a temperature gradient rather than a pressure gradient and a small pressure drop across the membrane is required to maintain the fluid/fluid interface at the mouth of the pore. Furthermore, the membrane imparts no selectivity to the separation (Gabelman and Hwang, 1999). Membrane contactors offer significant advantages over conventional dispersed-phase contactors such as the absence of emulsions, no density difference required between the fluids, easy scale-up, high interfacial area, no flooding at high flow rates. Within membrane contactors, membrane distillation (MD) and osmotic distillation (OD) seem to be a valid alternative for pressure-driven membrane processes in the concentration of liquid foods. In these processes, schematically represented in Fig. 9.1, the driving force for mass transfer is a vapour pressure difference across the membrane generated by either a temperature gradient (in MD) or a water activity difference (in OD). Their main advantages over traditional pressure-driven membrane processes are in terms of: low fouling, possibility of treatment of highly viscous solutions, high retention of species, and low energy consumption. Furthermore, these processes are not limited by high osmotic pressures and allow concentration levels to be attained that are similar to those obtained in thermal evaporation. 9.4.1 Osmotic distillation Process fundamentals Osmotic distillation is a recent membrane process also known as osmotic evaporation, osmotic concentration by membrane, membrane evaporation, isothermal membrane distillation, transmembrane distillation, thermopervaporation or gas membrane extraction. It can be carried out at room temperature and atmospheric pressure with minimal thermal and mechanical damage of the solutes. Consequently, it represents an attractive process for © Woodhead Publishing Limited, 2010
Advances in membrane-based concentration 251 Evaporation
Condensation
Tf > Tp Vapour
L/G interphase
Hydrophobic membrane
L/G interphase
Feed
Permeate (a)
Evaporation
Condensation
awf > aws Vapour
L/G interphase
Hydrophobic membrane
Feed
L/G interphase Stripping solution
(b)
Fig. 9.1 Schematic representation of (a) membrane distillation and (b) osmotic distillation. L/G, liquid–gas.
the concentration of solutions containing thermosensitive compounds, such as fruit juices and pharmaceuticals. In OD, a hydrophobic microporous membrane separates two liquid phases having different solute concentrations: a dilute solution on one side and a hypertonic salt solution on the opposite side. The hydrophobic nature of the membrane prevents penetration of the pores by aqueous solutions, creating air gaps within the membrane. The difference in solute concentration and, consequently, in water activity of both solutions, generates, at the vapour–liquid interface, a vapour pressure gradient across the membrane causing a vapour transfer across the pores from the high-vapour pressure phase to the low one (Alves and Coelhoso, 2002; Hogan et al., 1998; Lebfevre, 1988; Mengual et al., 1993). The concentration profile of the OD process is schematically presented in Fig. 9.2. The water transport through the membrane can be summarized in three steps: (1) evaporation of water at the dilute vapour–liquid interface; (2) diffusional vapour transport through the membrane pore; (3) condensation of water vapour at the membrane/brine interface (Schofield et al., 1987; Hogan et al., 1998). © Woodhead Publishing Limited, 2010
252 Separation, extraction and concentration processes High water vapour pressure Low water vapour pressure
Vapour flux
Feed dilutes aqueous solution
Csb Pwf >> Pwp Pore
Csm
Stripping solution (concentrated salt solution)
Cfm Cfb
Air Microporous hydrophobic membrane
Bulk Boundary layer
Bulk Boundary layer
Fig. 9.2 Concentration profile in osmotic distillation.
Because water transport involves condensation and evaporation phenomena, a temperature gradient through the membrane is generated, even if bulk temperatures of solutions separated by the membrane are equal. Consequently, a heat transfer tending to reduce the driving force for the water transport should be considered in addition to a mass transfer (Celere and Gostoli, 2002). The stripping solution after its dilution by water transferred from the feed stream can be reconcentrated by evaporation and reused in the OD operation. Therefore it must be thermally stable and preferably nontoxic, noncorrosive and available at low cost. A number of water-soluble salts such as NaCl, CaCl2, K2HPO4 are suitable. Salts displaying increase in solubility with temperature are preferred because they can be concentrated to high levels avoiding crystallization phenomena. Potassium salts of orthoand pyrophosphoric acid offer several advantages, including low equivalent weight, higher water solubility, steep positive temperature coefficients of solubility and safety in foods and pharmaceuticals (Deblay, 1995; Michaels, 1998). Typical transmembrane pressures encountered in osmotic distillation are in the range of 140 kPa (Hogan et al., 1998). Mass transfer aspects The water transport in the OD process is related to the driving force, represented by the vapour pressure difference at both liquid–vapour interfaces of the membrane (DPwm ), by the following equation:
Jw = Km (Pwf – Pws)
[9.1]
where Jw is the transmembrane flux, Km the membrane mass transfer coefficient, Pwf and Pws are the water vapour pressures of the feed and stripping solution at the membrane surface.
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Advances in membrane-based concentration 253 The water vapour pressures of the feed and stripping solutions at the fluid membrane interfaces are calculated as:
* Pwf = Pwf awf
[9.2]
* Pws = Pws aws
[9.3]
* * and Pws where Pwf are the pure water vapour pressures of the feed and stripping solution at the interface, respectively; awf and aws are the water activities of the feed and stripping solution at the interface, respectively. If the temperature is the same on both sides of the membrane equation [9.1] can be written as:
J w = K m Pw* (awf – aws )
[9.4]
The driving force depends on the solute concentration as well as on the temperature conditions prevailing at the vapour–liquid interfaces. A more detailed representation referring to the bulk conditions of both compartments is given by integrating the various mass transfer resistances:
lw = KD Pwb
[9.5]
where
1 1ˆ Ê1 K =Á + + ˜ K K K Ë f m s¯
–1
is defined as the overall mass transfer coefficient which accounts for the resistances opposed by the feed solution (1/Kf), the membrane (1/Km) and the stripping solution (1/Ks). If the concentration polarization is negligible, mass transfer depends only on membrane resistance and consequently:
K = K m
[9.6]
As in other membrane processes, OD is affected by concentration polarization. Owing to the water transport across the membrane, concentration in the bulk phase and, consequently, the water activity, differs from the composition near the membrane interface. As a result, the driving force for mass transfer and water flux are reduced (Ravindra Babu et al., 2006). Independently of the type of driving force (temperature difference, activity difference or both), the water vapour transfer mechanism in the membrane pores can be estimated on the basis of the molecular diffusion model, the Knudsen diffusion model or by a combination of both mechanisms (Geankoplis, 1993). If the mean free path of gas molecules is significantly greater than the membrane pore size, diffusing molecules collide frequently with pore
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254 Separation, extraction and concentration processes walls rather than with gas molecules present in the pores. In this instance the molecular transport occurs by Knudsen diffusion: 2e r K mK = M RT 3 cd
8RT pM
[9.7]
where r is the pore radius, e the membrane porosity, c the tortuosity factor, d the membrane thickness, M the molecular weight, R the gas constant and T the temperature. For pore size greater than the mean molecular free path of water vapour, the molecular diffusion is the controlling mechanism and the water mass flux can be written as: De P K mM = M RT cd Plm
[9.8]
where D is the diffusion coefficient and P the pressure. The Knudsen number (Kn) can be used to determine which of the two diffusion models is predominant. It compares the mean molecular free path (l) with the mean pore diameter of the membrane and is defined as: Kn = l 2p
[9.9]
where
l=
kBT P 2p s 2
[9.10]
in which kB is the Boltzmann constant, s is the mean collision diameter of the molecule and r the pore radius. For small pore size, Kn ≥ 10, collisions with pore walls are frequent therefore Knudsen diffusion is the prevailing mechanism. With relatively large pores, Kn ≤ 0.01, more collisions between the gas molecules themselves occur and the molecular diffusion will be predominant. Both mechanisms coexist when Kn is between the two limit values (Courel et al., 2000a). The concentration polarization phenomenon can be reduced by optimization of process parameters as stirring or stripping solution properties. The liquid concentration difference between the membrane and the bulk phase determines a diffusive solute flow according to Fick’s law which counterbalances the convective flow in the opposite direction. A mass balance across the boundary layer of the stripping solution is given by:
ÈÊ C ˆ ˘ J v = kp ln ÍÁ b ˜ ˙ ÎË Cm ¯ ˚
with
kp =
Dsw d
[9.11]
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Advances in membrane-based concentration 255 where Jv is the volume flux, Dsw the diffusion coefficient of solute by the membrane, kp the liquid mass transfer coefficient, Cm the salt molar concentration in the membrane, Cb the salt molar concentration in the bulk and d the boundary layer thickness. The liquid mass transfer coefficients depend on the solution physical properties and on the hydrodynamic properties of the system. These coefficients can be estimated by empirical correlations of dimensionless numbers namely Sherwood (Sh), Reynolds (Re) and Schmidt (Sc) numbers, equations [9.12] and [9.13].
Sh = a Reb Scg
[9.12]
with
Sh =
kdh , Dw
Re =
r udh , m
Sc =
m r Dw
[9.13]
where a, b and g are constants, k the feed or osmotic agent side liquid mass transfer coefficient, dh the hydraulic diameter, Dw the diffusion coefficient in water, r the solution density, m the solution dynamic viscosity and u the cross flow velocity of the fluid. Heat transfer OD is essentially a mass transfer process. The water transport through the membrane implies evaporation at the feed side and condensation at the stripping solution side. This phase change at the membrane walls generates a temperature difference across the membrane which reduces the vapour pressure difference decreasing the water transport. Under steady state conditions, the heat transfer equations can be written as:
Q = hf (Tf – Tfm) = hs (Tsm – Ts)
[9.14]
Q = Nw DHv – hm(Tsm – Tfm)
[9.15]
where Q is the heat flux, hf, hs and hm the heat transfer coefficients of the boundary layers (feed and stripping solution, respectively) and of the membrane, Tf and Tm the bulk temperatures of the feed and stripping solution, respectively, Tfm and Tsm the temperatures of the feed and stripping solution at the membrane interfaces, respectively, DHv the water latent heat of vaporization and Nw the molar vapour flux. The relation between heat flux and overall heat transfer coefficient during the OD process can be obtained by substituting Tsm and Tfm from equation [9.14] in equation [9.15], obtaining:
N DH Q = U ÈÍ w v – (Ts – Tf )˘˙ Î hm ˚
where U is the overall heat transfer coefficient given by: © Woodhead Publishing Limited, 2010
[9.16]
256 Separation, extraction and concentration processes
1 1ˆ Ê1 U=Á + + Ë hf hm hs ˜¯
–1
[9.17]
The thermal effect is considered negligible by most researchers. For example, Ravindra Babu et al. (2006, 2008) found that the contribution of concentration polarization on driving force reduction, during the concentration of pineapple and sweet-lime juices by OD, was prominent when compared with that of temperature polarization. On the other hand, the thermal effect owing to evaporation and condensation at both liquid–membrane interfaces was considered substantial by other authors (Courel et al., 2000a; Gostoli, 1999). In particular, the results of pure water OD experiments performed by Courel et al. (2000a) with commercial asymmetric porous membranes indicated that a high vapour flux of 12 kg m–2 h–1 generates a transmembrane temperature gradient of about 2 °C inducing a 30% driving force reduction. Osmotic distillation (OD) membranes and modules Hydrophobic polymers with low surface free energy are commonly used to produce membranes with pore sizes and pore size distribution suitable for OD applications. They include both polyolefins, such as polyethylene (PE) and polypropylene (PP), and perfluorocarbons such as polytetrafluoroethylene (PTFE) and polyvinylidene fluoride (PVDF). Membranes comprising these polymers, having a pore diameter ranging from 0.1 to 1 mm, allow the gas phase to be maintained in the membrane pores, a fundamental condition required to perform the OD process (Gryta, 2005). However, hydrophilic ceramic membranes can be modified by grafting on the surface molecules containing hydrophobic fluorocarbon chains like fluoroalkylsilanes or by coating the surface of alumina membranes with a thin lipid film. These membranes have been applied successfully in OD (Gabino et al., 2007; Romero et al. 2006). By referring to the configuration, hollow-fiber membranes, characterized by thin walls, are preferred for OD applications because they offer high surface/ volume ratios and do not require supports or spacers. Various parameters have to be considered in the selection of the membrane such as: pore size, porosity, conductivity and thickness. An increase in the pore size enhances the evaporation flux (flux is proportional to the radius); furthermore, membranes with large pore size exhibit a higher retention towards volatile organic flavour/ fragrance components (Barbe et al., 1998) than membranes with smaller pore size. These results can be attributed to differences in feed–membrane and stripper–membrane boundary layer resistances to organic volatiles transport owing to various degrees of liquid intrusion into the pores. Highly porous membranes are preferred because the evaporation flux is proportional to the porosity. Furthermore, an increase in membrane conductivity allows the heat gradient across the membrane to be minimized. Because the evaporation flux is proportional to the reciprocal of the pore length, the membrane thickness should be as low as possible. The thickness © Woodhead Publishing Limited, 2010
Advances in membrane-based concentration 257 is usually limited by the mechanical strength of the membrane: therefore thin membranes are supported by net, e.g. Gelman PTFE membranes. The overall thickness for OD membranes can vary from 80 to 250 mm, depending on the absence or presence of support. The risk of wetting of the hydrophobic membrane, with a consequent reduction in the evaporation flux and separation performance, is the main drawback of the osmotic distillation process. The reduction in the membrane thickness leads to a decrease in thermal resistance that facilitates the heat transfer from the brine to the feed causing a reduction in the temperature gradient (DT) and an increase in the partial pressure gradient (DP) across the membrane. However, thin hydrophobic membranes are more susceptible of wetting with a consequent loss of both evaporation flux and separation performance. If wetting occurs, the liquid can penetrate into pores in the membrane and, consequently, a liquid flux is added to the vapour flux and nonvolatile solutes diffuse across the membrane from one compartment to the other. Some fruit juices, especially citrus juices, contain peel oils and other highly hydrophobic compounds that promote wetting of OD membranes. Furthermore, cleaning solutions often contain surface-active agents that might also promote membrane wetting. Hydrophilic polymer films, which have sufficiently high intrinsic water permeability and are virtually impermeable to macrosolutes and colloids, are suitable as laminated membranes to prevent liquid intrusion without impeding vapour transport (Hogan et al., 1998). They include: esters and ethers of cellulose, crosslinked gelatin, chitin, agar, alginic acid, crosslinked polyacrylamide, crosslinked polyvinyl alcohol (PVA), polyhydroxy 2-ethyl methacrylate (PHEMA). Commercially available cellophane membranes can be suitable for this purpose; the hydrogel-film side of the laminate should be in contact with the solution to be concentrated (Michaels, 1999). Mansouri and Fane (1999) developed modified hydrophobic membranes for OD tolerant to oily feeds (i.e. limonene) by coating the feed side of commercial membranes with a thin layer of PVA. The uncoated membranes were rapidly wetted out by the oily feed even for low concentrations of oil (limonene) dispersed in water; vice versa, the coated membranes were stable for concentrations up to 1 wt.% limonene solution for periods up to 24 h. Commercial membranes for OD applications (termed TF200, TF450 and TF1000) manufactured by Pall-Gelman (East Hills, NY, USA) are characterized by a thin PTFE layer supported by a polypropylene (PP) net. In these asymmetric membranes, the top layer offers a resistance to gas transfer, whereas the membrane support offers an additional resistance to water transfer in the liquid form. The mass transfer resistance in the vapour phase is about 40–70% of the total resistance. The resistance of diluted brine entrapped in the PP support can cover up to 30% of the total resistance and the diluted brine boundary layer up to 60% indicating the sensitivity of the OD system to concentration polarization phenomena (Courel et al., 2000a). The best-known module designed for OD is the Liqui-Cel® Extra-Flow membrane contactor (Membrana-Charlotte, North Carolina, USA) containing
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258 Separation, extraction and concentration processes microporous PP hollow fibre membranes approximately 300 mm in external diameter with a wall thickness of about 40 mm; they have a mean pore diameter of about 30 nm and a porosity of about 40%. The fibres are potted into a polyethylene tubesheet and the shell casing is PP, PVDF or 316L stainless steel (Gabelman and Hwang, 1999; Hogan et al., 1998; Sirkar, 1997). The smallest modules are 2.5 ins in diameter with a membrane surface area of 1.4 m2, whereas the largest are 10 ins in diameter and offer a contact area of 130 m2. Commercial membranes commonly used in OD are summarized in Table 9.2. Effect of operating conditions on the OD flux The performance of the OD process in terms of water vapour transport is influenced profoundly by the operating conditions. Feed concentration inversely affected the performance of the OD process. Courel et al. (2000b) observed a decrease in evaporation flux (from 10.3 to 1.1 kg m–2 h–1) when sugar solutions of increasing sucrose content, from 0 to 65% w/w, were dehydrated at 25 °C by using stripping solution of 45.5 w/w% initial CaCl2 content. A transmembrane flux decay by increasing the feed concentration was observed also by Ravindra Babu et al. (2006) in the concentration of sweetlime juice and phycocianin solution by OD. The viscosity of the feed solution increases exponentially with the solute content whereas the diffusion coefficient strongly decreases relative to pure water. The increasing viscosity results in an increase of the concentration polarization effect reducing the driving force and, consequently, the flux rate (Bui et al., 2005). Sheng et al. (1991) studied the effect of operating conditions on the OD flux during the concentration of apple, orange and grape juice through a PTFE membrane with a pore size of 0.2 mm and an overall thickness of 100 mm. The osmotic pressure difference D p between the aqueous streams strongly affected the transmembrane vapour flux; in particular, a 33% decrease in D p determined a five-fold decline in OD flux. The OD flux is significantly affected also by the solute content of the stripping solution. Courel et al. (2000b) observed a 64% vapour flux decay for a 30% mass fraction reduction of the brine solution in extracting pure water at 25 °C by using CaCl2 as extractant. Ravindra Babu et al. (2006) observed a similar trend in the concentration of phycocianin solution and sweet-lime juice. These results may be explained by assuming the strong dependence of the water activity of the stripping solution on salt content. Although a vapour flux improvement is expected when the salt content is reduced (since the density and viscosity of the brine tend to decrease reducing the mass transfer resistance in the salt solution) this improvement is in fact masked by the activity effect. The OD flux is also differently affected by the type of osmotic agent. Calcium chloride produces higher transmembrane fluxes than sodium chloride © Woodhead Publishing Limited, 2010
Table 9.2 Typical membranes used in the OD process (adapted from Gryta, 2005) Manufacturer
Material
Configuration
Thickness (mm)
Porosity (%)
Average pore size (mm)
Durapore GVHP Durapore GVSP Durapore HVHP FHLP Accurel PP Q3/2 Accurel PP S6/2 Celgard 2500 Celgard 2400 Metricel SM35 TF200 TF450 TF1000 Gore-Tex 10387
Millipore Co. Millipore Co. Millipore Co. Millipore Co. Enka A.G. Enka A.G. Celgard LLC Celgard LLC Gelman USA Scimed Life Systems Inc. Pall-Gelman Pall-Gelman Pall-Gelman Gore & Associates
PVDF PVDF PVDF PTFE PP PP PP PP PP PDMS PTFE layer PTFE layer PTFE layer PTFE layer
Flat Flat Flat Flat Capillary Capillary Flat Flat Flat
125 108 125 175 200 400 25 25 90 300 178 178 178 8.5
70 80 75 70 70 70 50 37 55
0.20 0.22 0.45 0.25 0.2 0.2 0.07 0.05 0.1
80 80 80 78
0.2 0.45 1.0 0.2
supported supported supported supported
by by by by
PP PP PP PP
net net net net
Flat Flat Flat Flat
Advances in membrane-based concentration 259
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Type
260 Separation, extraction and concentration processes in the concentration of pineapple juice by OD (Nagaraj et al., 2006). This is mainly because of the higher osmotic activity of calcium chloride that results in a higher vapour pressure gradient across the membrane. Aqueous solutions of propylene glycol, glycerol and glycerol–salt mixtures were investigated as an alternative to calcium chloride in order to overcome the problem of corrosion and scaling associated with the use of brines (Celere and Gostoli, 2004). Propylene glycol and glycerol solutions (70–75 wt.%) were less effective than highly concentrated CaCl2 and exhibited a similar extractive power. However, propylene glycol cannot be recommended as an extractant in juice concentration owing to its low penetration pressure through the membrane pores and the not negligible volatility. Ternary mixtures water–glycerol–NaCl are characterized by lower viscosities than glycerol alone, and offer similar fluxes. The transmembrane flux in OD increases with an increase in the flow rate of the osmotic agent (Courel et al., 2000b; Nagaraj et al., 2006). This can be attributed to a stronger shear stress along the condensation side of the membrane leading to a reduction in the hydrodynamic boundary layer thickness and, consequently, to the concentration polarization effect. Ravindra Babu et al. (2008) observed an 8% increase in transmembrane flux when the flow rate of pineapple juice at about 12 °Brix was increased from 25 mL min–1 to 100 mL min–1; the increase in transmembrane flux can be explained by assuming a reduction in the concentration polarization effect on the feed side. The increase in flux, however, was more prominent (about 20%) by increasing the osmotic agent velocity. This phenomenon can be attributed to a lower concentration polarization on the feed side than on the brine side. A minor role of feed velocity on reducing the polarization problem in OD was observed also by Bui and Nguyen (2006) in the concentration of aqueous glucose solutions. They assumed that the heat transfer in OD is only a little involved or not involved at all; on the contrary, feed flow rate is essential to maintain the temperature gradient across the membrane in MD. The evaporation flux in OD is also affected by the feed temperature. Courel et al. (2000b) studied the influence of the temperature under isothermal conditions on the evaporation fluxes of pure water and sugar solutions of 35–65 w/w% sucrose content at 20–35 °C. A calcium chloride solution with an initial concentration of 45 w/w% was used as stripping solution. Evaporation fluxes ranged from 0.5 kg m–2 h–1 for a sugar solution of 65 w/w% (at 20 °C) to 12.8 kg m–2 h–1 for pure water (at 35 °C). The extent of mass transfer increase depended on the solute content: in the range 20–35 °C the evaporation flux increased by 120%, for a 35 w/w% sucrose solution, and only by 32% at 60 w/w% solute content. This phenomenon can be explained by assuming an exponential type relation between the vapour pressure difference across the membrane and the temperature according to Clapeyron’s law. Moreover, an increase in temperature results in a decrease in the feed and brine viscosities and an increase in the solute diffusion
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Advances in membrane-based concentration 261 coefficient: the higher the solute content, the stronger the thermal effect. Bui and Nguyen (2006) reported a 200% increase in the evaporation flux for a feed temperature increase of 20 degrees in the concentration of 40 w/w% and 50 w/w% aqueous glucose solutions by means of PVDF hollow fibres. OD applications Because OD can be operated at room temperature and atmospheric pressure, with minimal thermal and mechanical damage of the solutes, it has been successfully applied to the concentration of liquid foods such as milk, fruit and vegetable juices, instant coffee, tea and pharmaceuticals. The low temperature employed avoids the chemical or enzymatic reactions associated with heat treatment. The low operating pressure results in lower energy consumption and capital investment, reduced fouling phenomena and the possibility of using membranes characterized by lower mechanical resistance than those in pressure-driven membrane processes (Courel et al., 2001). OD is particularly suited for fruit juice concentration, producing concentrated juices with a quality and composition very close to fresh ones and higher than conventional products obtained by thermal evaporation (Cassano and Drioli, 2007; Jiao et al., 2004). Plate-and-frame modules with a net-shaped spacer on the extract side and a smooth juice side path were developed for the concentration of whole juice with a high pulp content (Cervellati et al., 1998). This configuration, even if not completely efficient in terms of mass transfer, allows unclarified juices to be processed. Helically wound hollow-fibre modules offer an improvement in the hydrodynamic conditions on the shell-side compared with axial flow modules; consequently, higher concentration of solutes and higher evaporation fluxes can be obtained when viscous feeds are processed (Costello et al., 1997). Rodrigues et al. (2004) evaluated the performance of the OD and RO processes in the concentration of camu-camu juice. RO allowed higher fluxes to be reached (50 kg m–2 h–1) than OD, but lower concentration of soluble solids (25 °Brix). OD allowed the juice to be concentrated up to 63 °Brix with evaporation flux values of 10 kg m–2 h–1. A pre-treatment of the whole juice by UF or MF allows evaporation fluxes to be improved in OD. However, the performance of the OD process is affected by pore diameters of UF membranes. Bailey et al. (2000) observed higher evaporation fluxes when Gordo grape juice was preliminarily ultrafiltered with membranes having pore diameters of 0.1 mm or less. UF membranes with a nominal pore diameter of 0.5 mm did not improve the OD fluxes. UF pretreatment results also in a small increase in juice surface tension with a consequent reduction in the tendency for membrane wet-out to occur. Several kinds of clarified fruit juices (orange, apple, grape, cactus pear, kiwi, passion fruit, and pineapple) were concentrated by OD (Alves and Coelhoso, 2006; Cassano and Drioli, 2007; Cassano et al., 2003, 2004, 2006,
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262 Separation, extraction and concentration processes 2007; Durham and Nguyen, 1994; Galaverna et al., 2008; Hongvaleerat et al., 2008; Jiao et al., 1991; Koroknai et al., 2006a, 2006b; Rektor et al., 2006; Rodrigues et al., 2004; Vaillant et al., 2001, 2005). An integrated membrane process was developed by Cisse et al. (2005) to produce concentrated orange juice. The process included a MF pre-treatment followed by OD. The quality of the concentrated juice was similar to the freshly squeezed one. Cassano et al. (2003, 2004, 2006 and 2007) proposed integrated membrane processes including UF, RO and OD to produce concentrated fruit juices such as kiwifruit, orange, lemon and cactus pear juice. Depectinized juices were clarified by UF and then optionally preconcentrated by RO. The OD process was employed as a concentration step to produce concentrated juices with final total soluble solids (TSS) up to 63–65 °Brix. The proposed process employs partial batch recycle on the feed side of the OD step in order to minimize feed viscosity changes through the membrane contactors and continuous counter current recycle plus evaporative reconcentration of the brine strip. The residual fibrous phase coming from the UF treatment (retentate) could be submitted to a stabilizing treatment (pasteurization, ohmic heating, high pressures, electric fields at high voltage) and successively added to the final OD concentrate for the preparation of fibre enriched beverages (Fig. 9.3). The total antioxidant activity (TAA) of the clarified or pre-concentrated juice was kept constant during the OD process independently of the achieved degree of TSS. Both aroma and colour were similar to those of the fresh juice. Ascorbic acid and health-promoting substances were also well preserved. On the contrary, a larger decrease in content of antioxidant compounds was detected in thermally evaporated juices with the same TSS content. In the concentrated orange juice, anthocyanins and hydroxycinnamates (particularly ferulic and p-coumaric acid) underwent a reduction of 36 and 55%, respectively; for the ascorbic acid and flavonones removals were in the order of 30% and 23%, respectively (Galaverna et al., 2008). Cassano and Drioli (2007) also evaluated the quality of kiwifruit juice concentrated by OD and thermal evaporation. The juice concentrated by thermal evaporation at 65 °Brix showed an 87% reduction of the ascorbic acid compared with the clarified juice. The TAA was reduced by about 50% independently of the TSS content achieved. On the contrary, the juice concentrated by OD retained almost all the TAA and the ascorbic acid content of the clarified juice. Blackcurrant juices concentrated by OD, with a final TSS content of 63–72 °Brix, showed a colour intensity, transparent ability and acidic flavour intensity similar to those of the raw juice (Kozák et al., 2008). A hybrid plant consisting of UF and RO pretreatment stages and an OD section (containing two 19.2 m2 Liqui-Cel membrane modules) for the concentration of fruit and vegetables juices was developed by Zenon Environmental (Burlington, Ont.). The plant was designed to produce concentrated juice at 65–70 °Brix at an average flow rate of 50 L h–1. It
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Fruit juice 10–11 °Brix
Reconcentrated brine
Feed reservoir Evaporator
Condenser
UF
Condensate
Stripping solution reservoir
OD Pulp
Preconcentrated juice 25–26 °Brix
RO
Water
Pasteurization
Pasteurized pulp
Reconstituted juice Concentrated juice 64–65 °Brix Water
Fig. 9.3 Integrated membrane process for the production of concentrated fruit juices (dashed lines refer to the production of the reconstituted juice).
Advances in membrane-based concentration 263
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Clarified juice
264 Separation, extraction and concentration processes was developed primarily for the concentration of grape juices destined for the production of high-quality vintage varietal wines such as Chardonnay, Cabernet Sauvignon and Merlot. Concentrates can be stored for long periods without deterioration and can be used as blending stocks to adjust the sugar content of freshly harvested grapes to minimize variations in alcohol content of the resulting wine between vintages. Additionally, the concentrated juice can be shipped over long distances and used to produce high-priced varietal wines starting from local grapes (Hogan et al., 1998). Evaporation fluxes between 7 and 10 kg m–2 h–1 were obtained by Hongvaleerat et al. (2008) in the concentration of clarified pineapple juice. These values were higher than those obtained with the single strength juice owing to the complete removal of suspended solids in the clarification step. At low TSS the evaporation flux decay during the OD process is more attributable to the dilution of the stripping solution whereas, at higher feed concentrations, the evaporation flux depends mainly on juice viscosity and, consequently, on juice concentration (Cassano and Drioli, 2007; Courel et al., 2000a; Vaillant et al., 2001). Evaporation fluxes in OD can be improved by application of acoustic fields. Narayan et al. (2002) observed an enhancement of the evaporation flux from 0.81 to 0.94 L m–2 h–1 during the OD of sugarcane juice with CaCl2 as stripping solution when ultrasounds were applied to the membrane cell. Concentration and temperature polarizations in OD contribute up to 18% to the flux reduction (Bui et al., 2005). However, concentration polarization gave a larger contribution to the flux reduction than temperature polarization. Furthermore, the flux reduction owing to polarization was smaller on the feed side than on the brine side. Hydrophobic membranes with relatively large pore sizes showed higher organic volatiles retention per unit water removal than those with smaller pores when model aroma aqueous solutions were processed by OD (Barbe et al., 1998). Pores with larger diameter at the membrane surface allow greater intrusion of the feed and brine solutions with a consequent increase in the thickness and resistance of the boundary layer at the pore entrance. Consequently, membranes with large surface pore diameters are preferred for OD applications in which retention of volatile flavour/fragrance components are required. Shaw et al. (2001) investigated the retention of flavours in concentrated orange and passion fruit juices prepared by OD. Losses of volatile components between 32 and 39% were observed in the concentrated products. The transfer of volatiles depends on the nature of compounds and on the operating conditions. A significant reduction in the transfer of juice volatiles was obtained by decreasing the circulation velocity or the temperature (Ali et al., 2002). The management of the diluted brine step is one of the drawbacks associated with the commercial application of the OD in fruit juice processing. Although the regeneration of exhausted brines can be realized by thermal evaporation, this operation is expensive owing to corrosion and scaling phenomena. Solar
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Advances in membrane-based concentration 265 ponding, reverse osmosis and pervaporation were proposed as alternatives to re-concentrate the diluted brine solutions (Thompson, 1991). Electrodialysis was also suggested for the regeneration of NaCl brines (Petrotos and Lazarides, 2001). The OD process is also a valid approach for producing low-alcohol beverages because it permits a selective removal of single volatile solutes from aqueous solutions (e.g. ethanol from wine and other ferments) by using water as a stripping solvent. The removal of organic solvents after the extraction of intracellular products (antibiotics, hormones and biologicals) from fermentation broths is a necessity in the drug industry. This operation has to be carried out at low temperature in order to preserve the product from thermal degradation and can be performed by OD in which water is employed as stripping solution (Hogan et al., 1998). The concentration of pharmaceuticals and biological compounds such as vaccines, hormones, recombinant proteins, enzymes, antibiotics, fungicides and nucleic acids is another potential use of OD. These products have to be isolated or recovered in the dry state in order to maintain their activity and shelf stability for a long time; OD can be used as a concentration step, eventually after a pre-concentration unit (by RO or NF), in order to obtain concentrates from which bioactive compounds can be more easily recovered (e.g. by crystallization or extraction). OD can be used also as a pre-concentration step for the lyophilization, reducing the water removal load during freeze drying. 9.4.2 Membrane distillation Process fundamentals Similarly to the OD process, in MD the water vapour transfer is promoted by a vapour pressure difference between two sides of a microporous hydrophobic membrane; however, in MD the physical origin of the vapour pressure difference is a temperature gradient rather than a concentration gradient: the feed is maintained at high temperature while cold water is used as a stripping permeate. Therefore, membrane distillation is a thermaldriven process. The mass transfer in MD can be described as a three-phase sequence: (1) formation of a vapour gap at the warm solution–membrane interface; (2) transport of the vapour phase through membrane pores; (3) its condensation at the cold side membrane–solution interface. Theoretical aspects and potential applications of the MD process have been extensively studied (Bandini et al., 1991; Calabrò et al., 1994; El-Bourawi et al., 2006; Lawson and Lloyd, 1997; Schofield et al., 1987; Tomaszewska, 2000a). The process takes place at atmospheric pressure and at temperatures that may be much lower than the boiling point of the treated solutions. Consequently, it can be used to concentrate solutes sensitive to high
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266 Separation, extraction and concentration processes temperatures such as fruit juices and pharmaceuticals. Feed temperatures in MD are typically 60–90 °C, although temperatures as low as 30 °C have been used. Operating pressures are generally of about 0–100 kPa, hence much lower than conventional pressure-driven membrane processes such as RO. Consequently, lower equipment costs and increased process safety can be achieved. Furthermore, the mechanical resistance of the membrane is greatly reduced. Because MD operates on the principles of vapour–liquid equilibrium, another advantage over traditional pressure-driven membrane processes is represented by its high rejection (theoretically 100%) towards ions, macromolecules, colloids, cells and other nonvolatile compounds (Lawson and Lloyd, 1997). Finally, vapour spaces can be reduced compared with conventional distillation processes. As in the OD process, the main drawback of MD is the risk of wetting of the hydrophobic membrane with a consequent reduction in the evaporation flux and separation performance. Therefore, the process solutions must be aqueous and sufficiently dilute. This limits MD to applications such as desalination, removal of trace volatile compounds from wastewater and concentration of nonvolatile aqueous solutions. Membrane distillation (MD) configurations As depicted in Figure 9.4, the MD process can be realized according to four types of configuration. In the direct contact membrane distillation (DCMD), the membrane separates the hot feed from the cold distillate. In this instance, the vapour pressure gradient, which results from the transmembrane temperature difference, is the driving force of the mass transport across the membrane. Volatile molecules evaporate at the hot liquid/vapour interface, cross the membrane pores in the vapour phase, and condense on the cold liquid/vapour interface inside the membrane module (Gryta, 2002; Lawson and Lloyd, 1997; Mengual and Peña, 1997; Tomaszewska, 2000b). DCMD is usually employed for applications in which water is the major fluxing component, such as desalination or concentration of aqueous solutions. In air gap membrane distillation (AGMD), a condensing surface is separated from the membrane by an air gap. Volatile molecules cross both the membrane pores and the air gap and finally condense over a cold surface inside the membrane module. In the sweeping gas membrane distillation (SGMD), a cold inert, gas sweeps the permeate side of the membrane carrying the volatile molecules. In this instance, condensation occurs outside the membrane module (Khayet et al., 2000). In the vacuum membrane distillation (VMD), vacuum is applied on the permeate side of the membrane by means of a vacuum pump and, similarly to SGMD, condensation takes place outside the membrane module (Bandini et al., 1992).
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Feed
Retentate
Feed
Retentate Membrane
Membrane
Permeate
Permeate in
Cooling water out
Air gap
Condensing plate
(a)
Feed
(b)
Retentate
Feed
Retentate Membrane
Membrane
Sweep gas in Condenser
Permeate
Sweep gas out
(c)
Cooling water in
Condenser
Permeate
Vacuum
(d)
Fig. 9.4 Membrane distillation configurations: (a) DCMD; (b) AGMD; (c) SGMD; (d) VMD.
Advances in membrane-based concentration 267
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Permeate out
268 Separation, extraction and concentration processes Mass transfer and polarization phenomena In a similar way to that of the OD process, classical models of gas diffusion in porous membranes, such as molecular diffusion or Knudsen diffusion, can be used to describe the mass transfer across MD membranes (Courel et al., 2000a; Gryta, 2005). These models suggest a linear relationship between the volume flux per unit surface area of the membrane and the transmembrane water vapour pressure difference at each interface, which depends on composition and temperature of both compartments. The osmotic and thermal contributions can operate either in a synergistic or in an antagonistic way (in which one of them prevails on the other) (Godino et al., 1995). The relationship between the permeate flux (Jw) and the driving force (DP) which causes the mass transfer across the membrane pores is given by:
Jw = Kw DP
[9.18]
where Km is the membrane mass transfer. The driving force in MD is a partial pressure gradient in the vapour phase given by:
DP = Pwf – Pwd
[9.19]
where Pwf and Pwd are water vapour pressures at the feed/membrane and feed/distillate interface, respectively. The mass transfer through the membrane, as in other membrane processes, is caused by the chemical potential difference (Dm) on both sides of the membrane (distillate and feed sides), which, for a ith component is given by:
D mi = D mid – D mif = RT ln
apid a = RT ln id apif ai
[9.20]
where apid and aid are the activity of ith component in the vapour and liquid if of ith component in phase on the distillate side, apif and aif the activity the vapour and liquid phase on the feed side, R the gas constant and T the temperature. The driving force expressed in equation [9.19] requires a knowledge of both temperature and solute composition at the vapour–liquid interfaces. Because the interfacial conditions are not always accessible, the water transport, referred to the bulk conditions, is often described as:
1 1ˆ Ê1 b b J w = K (Pwf – Pwd )=Á + + Ë K f K m K d ˜¯
–1
DP Pwb
[9.21]
where the overall mass transfer coefficient (K) is given by the series of water transport resistances from the bulk of the dilute solution towards the evaporation surface (1/Kf), water vapour transport through the membrane
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Advances in membrane-based concentration 269 (1/Km) and from the condensation surface to the bulk of the distillate side (1/Kd). The overall mass transfer coefficient depends on the membrane morphology, module design, solution concentration, temperature and hydrodynamic conditions (Gryta et al., 2005). When the temperatures in the layers near the membrane differ from those measured in the bulk of the feed and permeate side, a temperature polarization phenomenon occurs. However, temperature profiles formed in the DCMD are different from those observed in OD. The temperature gradient results from the evaporation at the feed side and condensation at the permeate side even if the bulk temperatures of two liquids are equal, as in OD. However, the heat conduction from the brine side to the feed side decreases the temperature polarization phenomenon in OD. On the contrary, in the case of DCMD, heat and mass transfer flow through the membrane in the same direction (from the feed to the distillate side) increasing the temperature polarization effect (Gryta, 2005). In all MD variants, as in OD, only water passes through the membrane causing an increase in the solute concentration on the feed side and a decrease in the solute concentration on the distillate side relative to the bulk condition. This phenomenon, called concentration polarization, results in a driving force reduction across the membrane. Several authors assume that the concentration polarization is substantial in OD, whereas the performance of DCMD is affected mainly by the temperature polarization effect, except for the concentration of solutions with a solute content close to the saturated state (Gryta, 2002). Alves and Coelhoso (2006) compared the performance of OD and MD in the concentration of sucrose solution (used as a model fruit juice), in terms of water flux and aroma retention. At the same applied overall driving force, water fluxes in MD were less than half of those observed in OD owing to temperature polarization effects. Moreover a higher retention of aroma compounds was observed for the OD process. MD membranes MD membranes must be porous and hydrophobic with good thermal stability and excellent chemical resistance to feed solutions. They must be characterized by high liquid entry pressure which is defined as the minimum hydrostatic pressure that must be applied to the feed liquid solution before it overcomes the hydrophobic forces and penetrates into the membrane pores. Membranes with high liquid entry pressure can be obtained by using materials with high hydrophobicity and small maximum pore size. The most suitable materials for hydrophobic MD membranes include PVDF, PP, PE and PTFE. These membranes are available in capillary or flatsheet configurations (Ding et al., 2002). PVDF appears of particular interest in MD processes because of its high melting point and good temperature resistance; its resistance to oxidation and gamma radiation, to solvents and to abrasion are also of interest. Hydrophilic membranes, such as cellulose
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270 Separation, extraction and concentration processes acetate membranes, can be treated in order to make their surfaces hydrophobic and, consequently, suitable for MD applications (Wu et al., 1992). Composite MD membranes consisting of hydrophobic and hydrophilic layers or a hydrophobic layer sandwiched between two hydrophilic layers have been also realized (Cheng et al., 1982). The size of micropores can range between 100 nm and 1.0 mm. In particular, the pores should be large enough to facilitate the required flux but also small enough in order to prevent liquid penetration through the membrane. For cylindrical pores, the maximum critical pore size at which the liquid penetrates the microporous phase is defined by the Laplace equation:
P = 2g
cos q r
[9.22]
where g is the surface tension of the liquid, q the contact angle between the liquid and the membrane, r the radius of the pore and P the applied pressure. For a given pore size a critical pressure Pc exists. For applied hydrostatic pressures higher than the Pc values, the liquid phase is transported across the membrane. Microporosity can be induced by mechanical stretching or by thermal phase separation technique. The relationship between the transmembrane flux and membrane characteristic parameters is given as:
N µ
ra e dm c
[9.23]
where N is the molar flux, r the mean pore size of the membrane pores, a a factor whose value equals 1 or 2 for Knudsen diffusion and viscous fluxes, respectively, dm the membrane thickness, e the membrane porosity and c the membrane tortuosity. According to equation [9.23], the MD flux increases with the increase in pore size. However, in order to avoid membrane pores wettability, the pore size should be as small as possible. Consequently, an optimum value for the pore size has to be determined for each MD application on the basis of the type of the feed solution to be treated. Schneider et al. (1988) recommended a maximum pore radius of 0.5–0.6 mm in order to avoid membrane wetting owing to fluctuations in process pressure and temperature. As for other membrane processes, the permeate flux in MD is inversely proportional to the membrane thickness: consequently thinner membranes produce higher fluxes. On the contrary, the membrane should be as thick as possible in order to prevent heat loss by conduction through the membrane matrix (Schofield et al., 1987). An optimal membrane thickness, considering the thermal conductivity of commercial membranes, should be within the range of 30–60 mm (Laganà et al., 2000). The porosity is the most influential factor affecting mass transfer rate in MD: membranes with higher porosity produce higher evaporation fluxes (Schneider et al., 1988). Membranes
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Advances in membrane-based concentration 271 with high porosity also allows the amount of heat lost by conduction to be reduced, because the heat transfer coefficient of the gas entrapped within the membrane pores is generally an order of magnitude smaller than that of the membrane material. MD membrane modules can be realized with flat-sheet membranes, as plate and frame and spiral-wound configurations, or with capillary membranes in tubular configuration. The design of the MD modules provides a high feed flow rate, to increase turbulence of the feed solution, low pressure drop and high packing density; moreover, because the MD is a nonisothermal process, good heat recovery and thermal stability is also guaranteed. Laboratory-scale modules are usually realized with flat-sheet membranes that are much more versatile than capillary membranes. They can be easily removed from their modules for cleaning, examination or replacement. Consequently, the same membrane module can be used to test different MD membranes. On the other hand, tubular membrane modules are preferred for commercial applications because they do not require a support and allow high membrane surface area/ module volume ratios to be realized. Effect of operating parameters on MD fluxes Evaporation fluxes in MD increase with an increase in the feed temperature. This is because the exponential increase of the vapour pressure of the feed solution with temperature increases the vapour pressure difference across the membrane and, hence, the driving force of the process. In DCMD applications the effect of decreasing the permeate temperature is to increase the evaporation flux (Lawson and Lloyd, 1999). A reduction in the flow rate implies a reduction in the Reynolds number and, consequently, in the transport coefficients. When the mean temperature is kept constant, the permeate flux increases linearly with the temperature difference. On the other hand, when the temperature difference is fixed, the permeate flux increases exponentially with the mean temperature (Mengual and Peña, 1997). When non-volatile solutes are considered, permeate fluxes in all MD configurations decrease by increasing the feed inlet concentration: this phenomenon can be attributed to the reduction of the driving force owing to the decrease in the vapour pressure of the feed solution and to the exponential increase in the viscosity of the feed solution. At high concentration ratios, MD fluxes are higher than those observed in other pressure-driven membrane processes (Cath et al., 2004). Permeate fluxes generally increase by increasing the solute concentration when aqueous solutions containing volatile components (such as alcohols) are processed. This can be explained by assuming that an increase in the volatile compound concentration in the feed side is associated with an increase in its transmembrane partial pressure. The increase in the permeate flow velocity causes an increase in the permeate flux in MD process. The increase in permeate flow velocity increases the heat transfer in the permeate side of the membrane reducing the temperature and concentration polarization phenomena. Consequently, the temperature
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272 Separation, extraction and concentration processes at the membrane surface approaches the temperature in the bulk permeate side increasing the driving force and the permeate flux. Finally, a linear increase in the MD flux with the transmembrane vapour pressure difference is typically observed in all MD configurations considered above. MD applications MD configurations previously described can be applied in different areas of interest for the separation of both nonvolatile (ions, colloids, macromolecules) and volatile (benzene, chloroform, trichloroethylene) compounds from water, the extraction of organic compounds such as alcohols from diluted aqueous solutions, the production of distilled water and the concentration of aqueous solutions. Desalination and pure water production from brackish water is the bestknown MD application (Banat and Simandl, 1998; Khayet et al., 2003, 2005). The concentration of radioactive solutions and wastewater treatment in the nuclear industry is another area under investigation (Khayet et al., 2006). In the chemical industry, the separation of azeotropic aqueous mixtures such as alcohol–water mixtures, the removal of volatile organic compounds from water and the concentration of acids, such as sulfuric, hydrochloric and nitric acid, can be achieved by MD (Banat and Simandl, 2000; Duan et al., 2001; Garcia-Payo et al., 2002; Tomaszewska, 2000b; Tomaszewska et al., 1995). Removal of dyes and wastewater treatment in the textile industry have also been reported (Banat et al., 2005; Calabrò et al., 1991). In the pharmaceutical and biomedical areas MD proved attractive for the removal of water from blood and protein solutions and in the wastewater treatment (Ortiz de Zárate et al., 1998; Sakai et al., 1986). MD can be successfully applied in the food industry and in areas where high temperature applications lead to the degradation of process fluids such as concentration of fruit juices and milk processing. A blackcurrant juice concentrate with a high TSS content was produced by MD without degradation of its valuable substances maintaining the juice side temperature at 26 °C and the water side temperature at 11 °C (DT = 15 °C) (Kozák et al., 2009). An increase of a few degrees Celcius in the driving force (DT = 19 °C) influenced significantly the evaporation flux and the operation time of the process. Gunko et al. (2006) used DCMD to investigate the concentration of apple juice. A TSS content of 50 °Brix was obtained when the permeate flux reached about 9 L m–2 h–1. Further concentration to 60–65 °Brix resulted in reduced productivity (up to 3 L m–2 h–1). Highly concentrated apple juices up to 64 °Brix were also produced by using polypropylene hollow fibre DCMD modules with trans-membrane fluxes of 1 kg m–2 h–1 (Curcio et al., 2000; Laganà et al., 2000). Flux rates were dependent essentially upon temperature polarization phenomena located mainly on the feed side, rather than concentration polarization. PVDF membranes used for the concentration of orange juice by MD
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Advances in membrane-based concentration 273 showed a very good retention of soluble solids, sugars and organic acids with rejection of sugars and organic acids equal to 100%. The colour and flavour of concentrated juice were satisfactory (Calabrò et al., 1994; Drioli et al., 1992). A protein concentrate containing at least 25% of protein was obtained by Christensen et al. (2006) in the processing of whey protein by DCMD as an alternative to thermal evaporation. The optimal temperature for the whey protein concentrate was 55 °C leading to less protein denaturation than evaporation and therefore a higher quality product. Bandini and Sarti (2002) studied the vacuum membrane distillation (VMD) for the concentration of grape must up to 50 °Brix. The process allowed production of juice concentrates that still retained interesting amounts of the aroma compounds. Bagger-Jørgensen et al. (2004) and Diban et al. (2009) evaluated the potential of VMD to recover aroma compounds from blackcurrant and pear juice, respectively. The highest values of enrichment factor (up to 15) for pear aroma compounds were obtained working at lower temperatures and higher downstream pressures. The highest concentration factors for the blackcurrant aroma compounds (from 21 to 31) were obtained at high feed flow rate (400 L h–1) and low temperatures (10 °C). At 5 vol.% feed volume reduction the recovery of highly volatile compounds was between 68 and 83 vol.% and between 32 and 38 vol.% for the hardly volatile compounds. Also in MD, an improvement in the evaporation flux in fruit juice concentration can be obtained when the juice is preliminarily submitted to a UF treatment. Drioli et al. (1992) found that the UF of the single-strength orange juice resulted in an increase in MD flux that remained almost constant during an approximately two-fold concentration. On the contrary, the MD flux of the unfiltered juice decreases steadily over the same concentration range. The improvement of the MD flux can be attributed to a reduction of juice viscosity as a result of pulp and pectin removal. PVDF membranes showed a very good retention of orange juice components such as total soluble solids, sugars and organic acids; on the other hand, the amount of ascorbic acid decreased by 42%, probably owing to its degradation associated with high temperature and oxidation. By referring to the retention of volatile organic flavour/fragrance components from liquid foods it has been found that membranes having an open fibrous structure rather than discrete pore offer the best volatiles retention for a given amount of removed water. 9.4.3 Coupled operation of osmotic distillation and membrane distillation Table 9.3 summarizes the main advantages and disadvantages of MD and OD processes. Because the water flux in OD is relatively low (usually between 0.07 and 7.2 kg m–2 h–1), the surface area available for flux should be increased. As described above, the thermal effect in OD reduces the driving
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274 Separation, extraction and concentration processes Table 9.3 Positive and negative aspects of MD and OD processes Advantages
Disadvantages
Low operating temperature
Low evaporative capacity with a long time of treatment Possibility of wetting of the hydrophobic membrane and consequent loss of flux and separation performance Production costs higher than those of thermal evaporation High cost of membrane replacement
Low operating pressure
No or less degradation of heat-sensitive compounds High TSS concentration in the concentrated products Modularity, easy scale-up Management of diluted brine solutions Possibility to treat solutions with high levels of suspended solids Possibility to concentrate several different products with the same unit No fouling problems Constant permeate flux in time Low investment cost
force of the water transport across the membrane. This phenomenon can be exploited to obtain a coupled process where the brine and feed solutions are thermostatically controlled at different temperatures: the osmotic solution on the cold side and the solution to be concentrated on the warm side. This coupled operation of MD and OD, referred to as membrane osmotic distillation (MOD), allows the water flux across the membrane to be enhanced (Wang et al., 2001). Although the aqueous solution is gently heated during the operation, this coupled method still works under mild conditions, because the temperature difference applied is lower than 15 °C (Bélafi-Bakó and Koroknai, 2006). The effect of varying experimental conditions (solute concentration, stirring rate, mean temperature and bulk temperature difference) on the water flux involved in a coupled process MD/OD was investigated by Godino et al. (1995). Koroknai et al. (2006a and 2006b) obtained concentrated fruit juices (apple, red- and blackcurrant, sour cherry and raspberry) at 60 °Brix in an operation time of 15–20 h, maintaining a temperature difference of 15 °C and using as an osmotic solution CaCl2 6M. In particular, the driving force of the process was greatly enhanced by decreasing the temperature of the osmotic solution, as in the range 15–25 °C only a slight drop in the water vapour pressure, for saturated CaCl2 solution, occurs. Consequently, maintaining the osmotic solution at room temperature, the energy consumption is minimized and the coupled process can be accelerated by increasing the temperature of the solution to be concentrated. Similarly, Bélafi-Bakó and Koroknai (2006) found that the coupled process is more effective than MD © Woodhead Publishing Limited, 2010
Process
Maximum achievable concentration (°Brix)
Product quality
Evaporation rate or flux
Possibility of treating Operating cost different products with the same installation
Thermal evaporation Cryoconcentration Reverse osmosis Direct osmosis Membrane distillation Osmotic distillation
65–75 30–50 25–30 50 60–70 60–70
Poor Very good Very good Good Good Very good
200–300 L h–1 – 5–10 L m–2 h–1 1–5 L m–2 h–1 1–10 L m–2 h–1 1–10 L m–2 h–1
No No No Yes Yes Yes
Moderate High High High High High
Capital investment
Energy consumption
Moderate Very high Moderate Moderate Low Low
Very high Very high High Low Low Low
Advances in membrane-based concentration 275
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Table 9.4 Key factors of conventional evaporation and membrane concentration techniques (adapted from Jariel et al., 1996)
276 Separation, extraction and concentration processes or OD alone. The use of short membrane modules in cascade series, with heat exchangers placed between them, was suggested in order to minimize heat losses.
9.5 Conclusions Key factors of conventional evaporation and membrane concentration techniques are summarized in Table 9.4. Several advantages of membranebased processes over conventional evaporation have been successfully demonstrated, including improved product quality, easy scale-up and low energy consumption. Low evaporation fluxes in OD and MD seem to be the main drawbacks of these processes when compared with RO and thermal evaporation. However, when the solution to be concentrated contain solutes sensitive to mechanical or thermal degradation there are serious limitations in using these technologies without significant deterioration in the quality. Currently membrane contactors are emerging technologies in the processing of foods, pharmaceuticals and beverages and will become breakthrough technologies when enhanced effectiveness is attained. Their integration with standard membrane operations is a valid approach for a sustainable industrial growth within the process intensification strategy. The aim of this strategy is to introduce in the productive cycles new technologies characterized by low encumbrance volume, advanced levels of automation capacity, modularity, remote control and reduced energy consumption. In order to improve the competitiveness of membrane contactors towards conventional technologies, efforts should be devoted to the development of new membranes characterized by high selectivity and stability for long-term applications, as well as to improvements in process engineering, including module and process design.
9.6 Nomenclature a a i api C d D h J k B k
water activity activity of the ith component in the liquid phase activity of the ith component in the vapour phase (fugacity) solute molar concentration (mol L–1) diameter (m) diffusion coefficient (m2 s–1) liquid heat transfer coefficient (W m–2 K–1) mass flux (kg m–2 h–1) or volume flux (m3 m–2 s–1) Boltzmann constant (1.3807 ¥ 10–23 J K–1) mass transfer coefficient (m s–1)
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Advances in membrane-based concentration 277 K Kn M N P P* Q r R Re Sc Sh T u U Greek g d e q DP Dm l m C r s
mass transfer coefficient (kg m–2 h–1 Pa–1) Knudsen number molecular weight (kg mol–1) molar flux (mol m–2 s–1) water vapour pressure (Pa) pure water vapour pressure (Pa) heat flux (W m–2) pore radius (m) gas constant (8314 J mol–1 K–1) Reynolds number Schmidt number Sherwood number temperature (°C, K) velocity of the fluid (m s–1) overall heat transfer coefficient (W m–2 K–1) symbols liquid surface tension (N m–1) thickness (m) porosity contact angle driving force chemical potential difference mean molecular free path (m) liquid dynamic viscosity (Pa s) tortuosity liquid density (kg m–3) mean collision diameter (m)
Subscripts b bulk d distillate f feed h hydraulic m membrane p pore or permeate s stripping solution or solute w water or vapour Superscripts b bulk location K Knudsen diffusion m membrane location M molecular diffusion
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278 Separation, extraction and concentration processes Abbreviations DO direct osmosis MD membrane distillation MF microfiltration OA osmotic agent OD osmotic distillation OMD osmotic membrane distillation PP polypropylene PTFE polytetrafluoroethylene PVDF polyvinylidene fluoride RO reverse osmosis TAA total antioxidant activity TSS total soluble solids UF ultrafiltration
9.7 References Aider M and de Halleux D (2008), ‘Production of concentrated cherry and apricot juice by cryoconcentration technology’, LWT-Food Sci Technol, 41, 1768–1775. Ali F, Dornier M, Duquenoy A and Reynes M (2002), ‘Transfer of volatiles through PTFE membrane during osmotic distillation’, Proceedings of the 2002 International Congress on Membrane and Membrane Processes, Toulouse, France. Alves VD and Coelhoso IM (2002), ‘Mass transfer in osmotic evaporation: effect of process parameters’, J Membrane Sci, 208, 171–179. Alves VD and Coelhoso IM (2006), ‘Orange juice concentration by osmotic evaporation and membrane distillation: a comparative study’, J Food Eng, 74, 125–133. Bagger-Jørgensen R, Meyer AS, Varming C and Jonsson G (2004), ‘Recovery of volatile aroma compounds from black currant juice by vacuum membrane distillation’, J Food Eng, 64, 23–31. Bailey AFG, Barbe AM, Hogan PA, Johnson RA and Sheng J (2000), ‘The effect of ultrafiltration on the subsequent concentration of grape juice by osmotic distillation’, J Membrane Sci, 164, 195–204. Banat FA and Simandl J (1998), ‘Desalination by membrane distillation: a parametric study’, Sep Sci Technol, 33, 201–226. Banat FA and Simandl J (2000), ‘Membrane distillation for propane removal from aqueous stream’, J Chem Technol Biotechnol, 75, 168–178. Banat FA, Al-Asheh S and Qtaishat M (2005), ‘Treatment of waters colored with methylene blue dye by vacuum membrane distillation’, Desalination, 174, 87–96. Bandini S, Gostoli C and Sarti GC (1991), ‘Role of mass and heat transfer in membrane distillation process’, Desalination, 81, 91–106. Bandini S, Gostoli C and Sarti GC (1992), ‘Separation efficiency in vacuum membrane distillation’, J Membrane Sci, 73, 39–52. Bandini S and Sarti GC (2002), ‘Concentration of must through vacuum membrane distillation’, Desalination, 149, 253–259. Barbe AM, Bartley JP, Jacobs AL and Johnson RA (1998), ‘Retention of volatile organic flavor/fragrance components in the concentration of liquid foods by osmotic distillation’, J Membrane Sci, 145, 67–75. Beaudry EG and Lampi KA (1990), ‘Membrane technology for direct osmosis concentration of fruit juices’, Food Technol, 44(6), 121.
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Advances in membrane-based concentration 279 Bélafi-Bakó K and Koroknai B (2006), ‘Enhanced water flux in fruit juice concentration: coupled operation of osmotic evaporation and membrane distillation’, J Membrane Sci, 269, 187–193. Bui AV, Nguyen HM and Muller J (2005), ‘Characterization of the polarizations in osmotic distillation of glucose solutions in hollow fibre module’, J Food Eng, 68, 391–402. Bui AV and Nguyen HM (2006), ‘The role of operating conditions in osmotic distillation and direct contact membrane distillation – a comparative study’, Int J Food Eng, 2(5) Art. 1, DOI: 10.2202/1556–3758.1171. Calabrò V, Drioli E and Matera F (1991), ‘Membrane distillation in the textile wastewater treatment’, Desalination, 83, 209–224. Calabrò V, Jiao B and Drioli E (1994), ‘Theoretical and experimental study on membrane distillation in the concentration of orange juice’, Ind Eng Chem Res, 33, 1803–1808. Cassano A, Conidi C, Timpone R, D’Avella M and Drioli E (2007), ‘A membrane-based process for the clarification and the concentration of the cactus pear juice’, J Food Eng, 80, 914–921. Cassano A, Drioli E, Galaverna G, Marchelli R, Di Silvestro G and Cagnasso P (2003), ‘Clarification and concentration of citrus and carrot juices by integrated membrane processes’, J Food Eng, 57, 153–163. Cassano A, Figoli A, Tagarelli A, Sindona G and Drioli E (2006), ‘Integrated membrane process for the production of highly nutritional kiwifruit juice’, Desalination, 189, 21–30. Cassano A, Jiao B and Drioli E (2004), ‘Production of concentrated kiwifruit juice by integrated membrane processes’, Food Res Int, 37, 139–148. Cassano A and Drioli E (2007), ‘Concentration of clarified kiwifruit juice by osmotic distillation’, J Food Eng, 79, 1397–1404. Cath TY, Adama VD and Childress AE (2004), ‘Experimental study of distillation using direct contact membrane distillation: a new approach to flux enhancement’, J Membrane Sci, 228, 5–16. Celere M and Gostoli C (2002), ‘The heat and mass transfer phenomena in osmotic membrane distillation’, Desalination, 147, 133–138. Celere M and Gostoli C (2004), ‘Osmotic distillation with propylene glycol, glycerol and glycerol–salt mixtures’, J Membrane Sci, 229, 159–170. Cervellati A, Zardi G and Gostoli C (1998), ‘Gas membrane extraction: a new technique for the production of high quality juices’, Fruit Process, 10, 417–421. Cheng DY and Wiersma SJ (1982), ‘Composite membrane for membrane distillation system’, US Patent 4,316,772. Cheryan M (1998), Ultrafiltration and microfiltration handbook, Technomic Publishing Co., Lancaster, PA. Christensen K, Andresen R, Tandskov I, Norddahl B and du Preez JH (2006), ‘Using direct contact membrane distillation for whey protein concentration’, Desalination, 200, 323–325. Cisse M, Vaillant F, Perez A, Dornier M and Reynes M (2005), ‘The quality of orange juice processed by coupling crossflow microfiltration and osmotic evaporation’, Int J Food Sci Technol, 40, 105–116. Collins AR and Harrington V (2002), ‘Antioxidants; not the only reason to eat fruit and vegetables’, Phytochem Revs, 1, 167–174. Costello AJ, Hogan PA and Fane AG (1997), ‘Performance of helically-wound hollow fibre modules and their application to isothermal membrane distillation’, Proceedings of Euromembrane ’97, 23–27 June, University of Twente, The Netherlands, 403–405. Courel M, Dornier M, Herry JM, Rios GM and Reynes M (2000b), ‘Effect of operating conditions on water transport during the concentration of sucrose solutions by osmotic distillation’, J Membrane Sci, 170, 281–289. Courel M, Dornier M, Rios GM and Reynes M (2000a), ‘Modelling of water transport
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280 Separation, extraction and concentration processes in osmotic distillation using asymmetric membrane’, J Membrane Sci, 173, 107– 122. Courel M, Tronel-Peyroz E, Rios GM, Dornier M and Reynes M (2001), ‘The problem of membrane characterization for the process of osmotic distillation’, Desalination, 140, 15–25. Curcio E, Barbieri G and Drioli E (2000), ‘Operazioni di distillazione a membrane nella concentrazione dei succhi di frutta’, Industria delle Bevande, XXIX, 113–121. Deblay P (1995), ‘Process for at least partial dehydration of an aqueous composition and devices for implementing the process’, US Patent 5,382,365. Diban N, Voinea OC, Urtiaga A and Ortiz I (2009), ‘Vacuum membrane distillation on the main pear aroma compound: experimental study and mass transfer modelling’, J Membrane Sci, 326, 64–75. Ding Z, Ma R and Fane AG (2002), ‘A new model for mass transfer in direct contact membrane distillation’, Desalination, 151, 217–227. Drioli E, Jiao B and Calabrò V (1992), ‘The preliminary study on the concentration of orange juice by membrane distillation’, Proceedings of VII International Citrus Congress, Acireale (Italy), 3, 1140–1144. Duan SH, Ito A and Ohkawa A (2001), ‘Removal of trichloroethylene from water by aeration, pervaporation and membrane distillation’, J Chem Eng Jpn, 34, 1069–1073. Durham RJ and Nguyen MH (1994), ‘Hydrophobic membrane evaluation and cleaning for osmotic distillation of tomato puree’, J Membrane Sci, 87, 181–189. El-Bourawi MS, Ding Z, Ma R and Khayet M (2006), ‘A framework for better understanding membrane distillation separation process’, J Membrane Sci, 285, 4–29. Gabelman A and Hwang S (1999), ‘Hollow fiber membrane contactors’, J Membrane Sci, 159, 61–106. Gabino F, Belleville MP, Preziosi-Belloy L, Dornier M and Sanchez J (2007), ‘Evaluation of the clearing of a new hydrophobic membrane for osmotic evaporation’, Sep Purif Technol, 55, 191–197. Galaverna G, Di Silvestro G, Cassano A, Sforza S, Dossena A, Drioli E and Marchelli R (2008), ‘A new integrated membrane process for the production of concentrated blood orange juice: effect on bioactive compounds and antioxidant activity’, Food Chem, 106, 1021–1030. Garcia-Payo MC, Rivier CA, Marison IW and Stockar UV (2002), ‘Separation of binary mixtures by thermostatic sweeping gas membrane distillation: II. Experimental results with aqueous formic acid solutions’, J Membrane Sci, 198, 197–210. Geankoplis CJ (1993), ‘Principles of mass transfer’, in Transport processes and unit operation, London, Prentice-Hall, 381–413. Godino MP, Peña L, Ortiz de Zárate JM and Mengual JI (1995), ‘Coupled phenomena membrane distillation and osmotic distillation through a porous hydrophobic membrane’, Sep Sci Technol, 30, 993–1011. Gostoli C (1999), ‘Thermal effects in osmotic distillation’, J Membrane Sci, 163, 75–91. Gryta M (2002), ‘Concentration of NaCl solution by membrane distillation integrated with crystallization’, Sep Sci Technol, 37, 3535–3558. Gryta M (2005), ‘Osmotic MD and other membrane distillation variants’, J Membrane Sci, 246, 145–156. Gunko S, Verbych S, Bryk M and Hilal N (2006), ‘Concentration of apple juice using direct contact membrane distillation’, Desalination, 190, 117–124. Herron JR, Beaudry EG, Jochums CE and Medina LE (1994), ‘Osmotic concentration apparatus and method for direct osmotic concentration of fruit juice’, US Patent 5,281,430. Ho WSW and Sirkar KK (1992), Membrane handbook, Chapman & Hall, New York. Hogan PA, Canning RP, Peterson PA, Johnson RA and Michaels AS (1998), ‘A new option: osmotic distillation’, Chem Eng Prog, 94, 49–61.
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Advances in membrane-based concentration 281 Hongvaleerat C, Cabral LMC, Dornier M, Reynes M and Nigsanond S (2008), ‘Concentration of pineapple juice by osmotic evaporation’, J Food Eng, 88, 548–552. Jariel O, Reynes M, Courel M, Durand N, Dornier M and Deblay P (1996), ‘Comparison of some fruit juice concentration techniques’, Fruits, 51, 437–450. Jiao B, Cassano A and Drioli E (2004), ‘Recent advances on membrane processes for the concentration of fruit juices: a review’, J Food Eng, 63, 303–324. Jiao B, Molinari R, Calabrò V and Drioli E (1991), ‘Application of membrane operations in concentrated citru juice processing’, Agro-Ind Hi-Tech, 19–27. Karode SK, Kulkarni SS and Ghorapade MS (2000), ‘Osmotic dehydration coupled reverse osmosis concentration: steady-state model and assessment’, J Membrane Sci, 164, 277–288. Khayet M, Godino MP and Mengual JI (2000), ‘Theory and experiments on sweeping gas membrane distillation’, J Membrane Sci, 165, 261–272. Khayet M, Mengual JI and Matsura T (2005), ‘Porous hydrophobic/hydrophilic composite membranes. Application in desalination using direct contact membrane distillation’, J Membrane Sci, 252, 101–113. Khayet M, Mengual JI and Zakrewska-Trznadel G (2003), ‘Theoretical and experimental studies on desalination using the sweeping gas membrane distillation’, Desalination, 157, 297–305. Khayet M, Mengual JI and Zakrewska-Trznadel G (2006), ‘Direct contact membrane distillation for nuclear desalination. Part II. Experiments with radioactive solutions’, Int J Nuclear Desalination, 56, 56–73. Koroknai B, Gubicza L and Bélafi-Bakó K (2006a), ‘Coupled membrane process applied to fruit juice concentration’, Chem Pap, 60, 399–403. Koroknai B, Kiss K, Gubicza L and Bélafi-Bakó K (2006b), ‘Coupled operation of membrane distillation and osmotic evaporation in fruit juice concentration’, Desalination, 200, 526–527. Kozák A, Bánvölgyi S, Vincze I, Kiss I, Békássy-Molnár E and Vatai G (2008), ‘Comparison of integrated large scale and laboratory scale membrane processes for the production of black currant juice concentrate’, Chem Eng Proc, 47, 1171–1177. Kozák A, Békássy-Molnár E and Vatai G (2009), ‘Production of black-currant juice concentrate by using membrane distillation’, Desalination, 241, 309–314. Laganà F, Barbieri G and Drioli E (2000), ‘Direct contact membrane distillation: modelling and concentration experiments’, J Membrane Sci, 166, 1–11. Lawson KW and Lloyd DR (1997), ‘Membrane distillation’, J Membrane Sci, 124, 1–25. Lebfevre MSM (1988), ‘Method of performing osmotic distillation’, US Patent 4,781,837. Luh BS, Feinberg B, Chung JI and Woodroof JG (1986), ‘Freezing fruits’ in Woodroof JG and Luh BS, Commercial fruit processing, Westport, AVI Publishing Co., 263–351. Maccarone E, Campisi S, Cataldi Lupo MC, Fallico B and Nicolosi Asmundo C (1996), ‘Thermal treatments effects on the red orange juice constituents’, Ind Bevande, 25, 335–341. Mansouri J and Fane AG (1999), ‘Osmotic distillation of oily feeds’, J Membrane Sci, 153, 103–120. Mengual JI, Ortiz de Zárate JM, Peña L and Velázquez A (1993), ‘Osmotic distillation through macroporous hydrophobic membranes’, J Membrane Sci, 82, 129–140. Mengual JI and Peña L (1997), ‘Membrane distillation’, Colloid Interf Sci, 1, 17–29. Michaels AS (1998), ‘Methods and apparatus for osmotic distillation’, US Patent 5,824,223. Michaels AS (1999), ‘Osmotic distillation process using a membrane laminate’, US Patent 5,938,928. Milleville H (1990), ‘Direct osmotic concentrates juices at low temperature’, Food Proc, 51, 70–71. © Woodhead Publishing Limited, 2010
282 Separation, extraction and concentration processes Mulder M (1998), Basic principles of membrane technology, 2nd edn. London, Kluwer Academic Publishers, 280–303. Nagaraj N, Patil G, Babu BR, Hebbar UH, Raghavarao KSMS and Nene S (2006), ‘Mass transfer in osmotic membrane distillation’, J Membrane Sci, 268, 48–56. Narayan AV, Nagaraj N, Hebbar HU, Chakkaravarthi A and Raghavarao KSMS (2002), ‘Acoustic field-assisted osmotic membrane distillation’, Desalination, 147, 149–156. Ortiz de Zárate JM, Rincón C and Mengual JI (1998), ‘Concentration of bovine serum albumin aqueous solutions by membrane distillation’, Sep Sci Technol, 33, 283–296. Petrotos KB, Quantick PC and Petropakis H (1998), ‘A study of the direct osmotic concentration of tomato juice in tubular membrane module configuration. I. The effect of certain basic process parameters on the process performance’, J Membrane Sci, 150, 99–110. Petrotos KB, Quantick PC and Petropakis H (1999), ‘Direct osmotic concentration of tomato juice in tubular membrane module configuration. II. The effect of using clarified tomato juice on the process performance’, J Membrane Sci, 160, 171–177. Petrotos KB and Lazarides HN (2001), ‘Osmotic concentration of liquid foods’, J Food Eng, 49, 201–206. Ravindra Babu B, Rastogi NK and Raghavarao KSMS (2006), ‘Mass transfer in osmotic membrane distillation of phycocyanin colorant and sweet-lime juice’, J Membrane Sci, 272, 58–69. Ravindra Babu B, Rastogi NK and Raghavarao KSMS (2008), ‘Concentration and temperature polarization effects during osmotic distillation’, J Membrane Sci, 322, 146–153. Rektor A, Vatai G and Békássy-Molnár E (2006), ‘Multi-step membrane processes for the concentration of grape juice’, Desalination, 191, 446–453. Rodrigues RB, Menezes HC, Cabral LMC, Dornier M, Rios GM and Reynes M (2004), ‘Evaluation of reverse osmosis and osmotic evaporation to concentrate camu-camu juice (Myrciaria dubia)’, J Food Eng, 63, 97–102. Romero J, Draga H, Belleville MP, Sanchez J, Come-James C, Dornier M and Rios GM (2006), ‘New hydrophobic membranes for contactor processes – applications to isothermal concentration of solutions’, Desalination, 193, 280–285. Sakai K, Muroi T, Ozawa K, Takesawa S, Tamura M and Nakaue T (1986), ‘Extraction of solute-free water from blood by membrane distillation’, Trans Am Soc Artif Intern Organs, 32, 397–400. Schneider K, Holz W and Wollbeck R (1988), ‘Membranes and modules for transmembrane distillation’, J Membrane Sci, 39, 25–42. Schofield RW, Fane AG and Fell CJD (1987), ‘Heat and mass transfer in membrane distillation’, J Membrane Sci, 33, 299–313. Shaw PE, Lebrun M, Dornier M, Ducamp MN, Courel M and Reynes M (2001), ‘Evaluation of concentrated orange and passion fruit juices prepared by osmotic evaporation’, Lebensm Wiss Technol, 34, 60–65. Sheng J, Johnson RA and Lefebvre MS (1991), ‘Mass and heat transfer mechanism in the osmotic distillation process’, Desalination, 80, 113–121. Sirkar KK (1997), ‘Membrane separation technologies: current developments’, Chem Eng Commun, 157, 145–184. Thompson D (1991), ‘The application of osmotic distillation for the wine industry’, Aust Grapegrower Winemaker, 11, 11–14. Tomaszewska M (2000a), ‘Concentration and purification of fluosilicic acid by membrane distillation’, Ind Eng Chem Res, 39, 3028–3041. Tomaszewska M (2000b), ‘Membrane distillation – examples of applications in technology and environmental protection’, Pol J Environ Stud, 9, 27–36. Tomaszewska M, Gryta M and Morawski AW (1995), ‘Study on the concentration of acids by membrane distillation’, J Membrane Sci, 78, 277–282. © Woodhead Publishing Limited, 2010
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284 Separation, extraction and concentration processes
10 Separation of value-added bioproducts by colloidal gas aphrons (CGA) flotation and applications in the recovery of value-added food products P. Jauregi and M. Dermiki, The University of Reading, UK Abstract: This chapter describes the application of a flotation method based on the formation of colloidal gas aphrons (CGA), which are surfactant-based microbubbles, to the recovery of value-added food products from fermentation broths and waste streams of food and agricultural industries. Their main properties are described and contrasted with those of conventional foams. The fundamentals of the separation are described and illustrated by several case studies on CGA applications to the recovery of whey proteins, polyphenols and carotenoids. Finally, a critical evaluation of the feasibility of CGA for industrial applications is carried out based on the astaxanthin case study and future applications of CGA are discussed. Key words: colloidal gas aphrons, flotation, bioproducts, surfactants, food ingredients, recovery.
10.1 Introduction Many studies have been conducted in recent years into additional biological functionalities of several food components and ingredients such as the dairy proteins, peptides, polyphenols or carotenoids so that they could be produced as ingredients in functional foods. The development of functional foods poses new challenges to the food industry as new processes and technologies are required. The food industry must adapt to these changing times and to the revolution caused by newly emerging functional foods. In addition, in the last few years reducing the environmental impact of industrial wastes has been a subject of increasing concern. In industrial wastewaters, these compounds considerably increase biochemical and oxygen demands, with detrimental effects on the flora and fauna of discharge zones, whereas in solid residues © Woodhead Publishing Limited, 2010
Separation of value-added bioproducts by colloidal gas aphrons 285 used as fertiliser they may inhibit germination. Furthermore sustainability is given high priority by governments and industries are encouraged to develop practices for the usage and recycling of wastes. Food processing waste streams and agricultural wastes can be exploited as cheap sources of high-value products. For example, proteins and peptides can be extracted from whey which is a by-product in cheese production and polyphenols can be extracted from winemaking waste, such as grape skin and marc. Separation and purification of these products is carried out by a number of steps including liquid–liquid extraction, membrane techniques and/or chromatographic methods. Alternative separations such as flotation may be used for the recovery stage although further processing will be necessary if high purity products are required. Foam fractionation has been applied for the recovery of proteins. Flotation, which is the foam separation of insoluble compounds, has been applied mainly to wastewater treatment and to the recovery of minerals. Separation of biomass has also been successfully carried out by flotation. This chapter deals mainly with a particular type of flotation using microbubbles which are also called colloidal gas aphrons (CGA). CGA were first described by Sebba (1972) as surfactant-stabilised microbubbles (10–100 mm) generated by intense stirring (>8000 rpm) of a surfactant solution. In this chapter, the application of CGA to bioseparations is reviewed particularly in relation to the recovery of value-added food products. First, the structural and dispersion characteristics of CGA are described. Then, the separation fundamentals are described and illustrated with the application of CGA to the separation of proteins. In addition, the feasibility of CGA for industrial applications is considered, focusing on: (i) scalability using a flotation column and (ii) removal and recycling of surfactant. All these aspects are illustrated in several case studies of CGA separations for the recovery of bioactives from plant extracts (norbixin and polyphenols) and recovery of astaxanthin produced by microbial fermentation.
10.2 Colloidal gas aphrons (CGA) properties 10.2.1 Structure of CGA Colloidal gas aphrons (CGA) are surfactant-stabilised microbubbles (10– 100 mm) generated by intense stirring of a surfactant solution at high speeds (>8000 rpm). Sebba (1987) postulated that they are composed of multilayers of surfactant molecules; as depicted in Fig. 10.1 surfactant molecules adsorb at the interface with hydrophilic heads towards the aqueous phase and the hydrophobic tails towards the gas phase. Sebba’s hypothesis on the structure was based on several experimental observations such as delayed coalescence and hence higher stability of CGA than conventional foams. Jauregi et al. (2000) investigated the drainage kinetics of CGA and compared measured drainage rates with those obtained by applying predictive models for foams with and without modifications. The main modifications were in relation to © Woodhead Publishing Limited, 2010
286 Separation, extraction and concentration processes Inner surface of shell
Air core
Viscous water Shell
Electrical double layer
Outer surface of shell
Fig. 10.1 Proposed structure of CGA by Sebba (1987).
structural differences between foams and CGA upon drainage, such as aphrons being surrounded by a liquid film so that they do not adopt a dodecahedral shape. Interestingly, the model giving the best prediction was the one including differences in structural features which further supports Sebbas’s theory. Jauregi and coworkers used for the first time small angle x-ray diffraction in an attempt to determine the thickness of the surfactant film and the number of surfactant layers of CGA generated by the anionic surfactant sodium bis(2-ethyl hexyl) sulfosuccinate (AOT). The analysis of the data proved to be difficult as for the same sample different film thickness values could be obtained. However, an interesting finding was that samples containing aphrons gave similar scattering regardless of the surfactant concentration and these corresponded to 5.4 nm (which is equivalent to seven layers of surfactant assuming the full length of the surfactant molecule arranged in layers and vertically at the interface). Moreover, the same surfactant solutions with no aphrons gave a different scattering signal and, in this instance, differences between different concentrations of surfactant solutions were found; for example, at concentrations around the cmc (critical micellar concentration), the scattering corresponded to a bilayer, which could correspond to micelles whereas, above the cmc, the scattering corresponded to a multilayer of 3–5 molecules, confirming the hypothesis that AOT forms lamellar structures (Jauregi et al., 2000). In a number of research studies predictive models for liquid drainage in CGA were developed based on measurements of liquid drainage rate (Table 10.1). Most of the studies highlight the difference between CGA and conventional foams. Bhatia et al. (2005) studied the effect of stirring speed and stirring time on the stability of CGA generated using a range of
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Table 10.1 Summary of studies for the characterisation of CGA (c, cationic; a, anionic; n-i, nonionic) (adapted from Dermiki, 2009) Surfactant
Parameters studied
Main findings
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Zhao et al., SDS (a) 2009; Larmignat CTAB (c) et al., 2008 Tween 20 (n-i)
Surfactant type, Csurf, pressure CGA rheology is not affected by pipe shape or hydraulic diameter, CGA drop, T, volumetric and mass can be treated as shear thinning fluid for all surfactant solutions. Increase flow rate, pipe diameters and in Csurf leads to increased shear stress for a given shear rate shape
Feng et al., 2009 Tergitol (n-i) Ramnolipid (a)
Csurf, pH, Csalt
Stability increases with Csurf and decreases with increasing pH, Csalt. Three distinct phases in the liquid drainage
Moshkelani and SDS (a) Amiri, 2008
Electrical conductivity of surfactant solution
Three clear stages of drainage of CGA can be identified, compared with two stages for conventional foams
Tseng et al., 2006
Pressure drop, T, mass and volumetric flow rate of CGA
CGA can be considered shear thinning fluid. Heat transfer coefficients for CGA made of water and surfactant are smaller than that of water, and they are constant and independent of mass flow and heat flux
tstir, rpm, Csurf, mixing two oppositely charged CGA
Development of an empirical correlation for variation of gas hold-up with tstir, rpm Ê t ˆ – 4.025ˆ a = (0.004 + 44.017) Á 0.2407 ln ÊÁ –2.00 ˜¯ ˜¯ Ë2 Ë
Bhatia et al., 2005
SLS (a) CTAB (c)
The mixing of the two oppositely charged CGA showed no effect on drainage Yan et al., 2005 SDS (a) CTAB (c) Tween 80 (n-i)
Type of surfactant, Csurf, T
There was no effect of the type of surfactant on the stability. Mathematical rn model which describes drainage of CGA: Vt = Vmax n . Two distinct K + rn stages of drainage are determined by two independent mechanisms
Jarudilokkul et al., 2004
Tween 20 (n-i)
tstir, pH, rpm, CNaCl, Csurf, protein separation
Increasing Csurf, tstir increases the stability Increase of Csurf leads to decrease of protein separation
Dai and Deng, 2003
HTAC (c)
pH, concentration of silica sol for stabilisation
CGA were stable for up to 12 h at concentration of silica sol 0.15–0.25 mol dm–3 and pH 7–10
Separation of value-added bioproducts by colloidal gas aphrons 287
Reference
© Woodhead Publishing Limited, 2010
Reference
Surfactant
Parameters studied
Main findings
Jauregi et al., 2000
AOT (a)
Csurf, pH, Csalt, T, tstir
X-ray diffraction for the characterisation of CGA showed evidence of the existence of surfactant multilayers
Jauregi et al., 1997
AOT (a)
Csurf, pH, salt, T, tstir
Increase of Csurf increases the stability, whereas addition of salt leads to the opposite effect Significant interactive effects: (i) Csurf · salt, (ii) pH · Csurf, (iii)T · tstir, (iv) Csurf · T Gas hold-up depends on tstir, Csurf, and salt
Bredwell et al., 1995
SDS (a) Csurf, salt CPC (c) Triton X-100 (n-i)
Save and DTAC, CTAC, Pangarkar, 1994 CPC, DMDSAC (c) SDBS, SLS (a) Chaphalkar et al., 1993
pH, tstir, Csurf, Csalt, viscosity, additives, impeller clearance
SDBS (a), CTAB Type and concentration of (c), surfactant, ionic strength Tergitol (n-i)
Amiri and TTAB (c) Woodburn, 1990
Csurf, pH
Increasing Csurf leads to decrease of formation time of CGA and increase in CGA stability. No effect of salt. No effect of surfactant type and surfactant concentration on the gas hold-up pH and impeller clearance had no effect on stability. Type of surfactant, viscosity, addition of enzymes, polymers, nonionic surfactants, solvents and salts affect gas hold-up and stability MDtergitol<MDCTAB<MDDDBS. Increase of Csurf leads to decrease of MD. Addition of salts reduces the MD of ionic surfactants There is agreement between the predicted and observed rise velocities Stability of CGA depends on the Csurf and pH
MD, mean diameter; tstir, stirring time; Csurf, concentration of surfactant; Csalt, concentration of salt; T, temperature; rpm, stirring speed. SDS, sodium dodecyl sulfate; SDBS, sodium dodecylbenzenesulfonate; CTAB, cetyltrimethyl ammonium bromide; SLS, sodium lauryl sulfate; AOT, sodium bis (2-ethyl hexyl) sulfosuccinate. HTAC, hecadecyl trimethyl ammonium chloride; CPC, cetyl piridinium chloride; DTAC, dodecyltrimethylammonium chloride; CTAC, cetyltrimethylammonium chloride; DMDSAC, dimethyl distearyl ammonium chloride; TTAB, tetradecyl trimethyl ammonium bromide.
288 Separation, extraction and concentration processes
Table 10.1 Continued
Separation of value-added bioproducts by colloidal gas aphrons 289 surfactants and they developed a first order kinetics model for the drainage. Interestingly, drainage kinetics did not change when they used a mixture of an anionic and a cationic surfactant as opposed to conventional foams (Bhatia et al,. 2005). Yan et al (2005) developed a mathematical kinetic model by studying CGA generated using a range of concentrations of surfactants at varying temperatures. According to this model there were two distinct stages of CGA drainage (Yan et al., 2005). On the other hand, in a more recent study Moshkelani and Amiri (2008) measured the electrical conductivity of CGA dispersions and found that there are three separate stages in the liquid drainage of aphrons as opposed to the two stages in conventional foams (Moshkelani and Amiri, 2008). This was further supported in a more recent study by Feng et al. (2009) where the volume of drained liquid was measured and also photographs of the foams at different stages were taken. At the first stage they showed how drainage increases with time owing to a combination of upflow migration of bubbles and downward liquid drainage under gravity. Then, in the following stage, the dispersion behaves like conventional foams and, at the final stage, the foam becomes water deficient and begins to behave like dry foam. At this stage, the foam drainage is very slow owing to slow liquid release from films under capillary suction , i.e.: liquid is released from the film owing to lower pressure in the capillaries between the bubbles (Feng et al., 2009). Overall these studies support further the hypothesis that there are structural differences between CGA and conventional foams. 10.2.2 Characteristics of CGA Owing to their unique structure CGA possess the following important properties: ∑
∑
Higher stability than conventional foams. Owing to the multilayer structure, CGA exhibit high stability. When two aphrons collide, the momentum may not be enough to break the barrier of six surfactantstabilised interfaces, as opposed to the two surfactant-interfaces when two bubbles of conventional foams collide (Sebba, 1987). Stability is measured in terms of the liquid drainage rate, which is generally considered to follow first order kinetics (Bhatia et al., 2005) consequently, half-life (t) is defined as the time required for half the original volume of liquid to drain. CGA are also characterised in terms of gas hold-up (e), which is defined as the ratio between the gas volume (Vg) and the dispersion final volume after stopping stirring and at time = 0 of drainage (Va0): e = Vg/Va0
[10.1]
The buoyancy of the encapsulated gas leads to easy separation of the aphron phase from the bulk liquid phase. Therefore, no centrifuges are needed to separate the two phases. On the other hand, creaming can be
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290 Separation, extraction and concentration processes avoided when necessary by stirring CGA at such a rate that the lateral movement conveyed to the bubbles is greater than the upward buoyancy owing to gravity. ∑ Adherence of particles and molecules to the encapsulating shell. It is possible to modify the surface properties of CGA by using different types of surfactant. Depending on the surfactant used to generate the CGA, the outer surface of the bubble can be negatively (anionic surfactant), positively (cationic surfactant) or non-charged (nonionic surfactant). Consequently, oppositely charged molecules adsorb. Thus, the selectivity of adsorption can be modified by changing the type of surfactant. This is illustrated in experiments carried out in our laboratory with two oppositely charged dyes, the cationic methylene blue, and the anionic methyl orange and CGA generated by the cationic surfactant CTAB (see Plate I, between pages 292 and 293). CGA are contacted with the aqueous solution containing both dyes and then the dispersion is allowed to settle so it separates into two phases (Plate Ic). The top CGA phase contains mainly methyl orange (orange colour top phase) and the liquid phase contains mainly methylene blue (blue colour liquid phase). Basu and Malpani (2001) also reported the effective separation of these dyes using CGA generated by CTAB. ∑ Low viscosity of the system. Flow properties are similar to those of water, as long as the gas hold-up does not exceed 65% (Roy et al., 1995), regardless of the type of the surfactant used to generate them. Consequently, CGA can be pumped easily from the generation point to the point where they are going to be used, whereas, for conventional foams, their characteristics change during pumping owing to the elastic nature of the bubble. This is advantageous when applied to a flotation column for fractionation and flotation in the removal and/or recovery of products. Several researchers have investigated the rheology of CGA, as shown in Table 10.1 (Larmignat et al., 2008; Tseng et al., 2006; Zhao et al., 2009). An interesting application which takes advantage of CGA rheological properties is described in a patent by Brookey (2004) where CGA are used to produce fluids with improved shear thinning properties, and this has led to a successful industrial application for well-drilling fluids by MASI technologies LLC. ∑ Small size of bubbles, resulting in larger interfacial area per unit volume, and thus a large capacity for adsorption of molecules and fine particles. Bubble size depends on the concentration and type of surfactant, ionic strength and the presence of other molecules or particles. Jauregi (1997) found that bubble size of AOT-generated CGA increased with surfactant concentration, ionic strength and stirring time up to 10 min. Size distribution of CGA generated by selected surfactants can be seen in Table 10.2.
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Separation of value-added bioproducts by colloidal gas aphrons 291 Table 10.2 Characteristics of CGA generated by ionic and nonionic surfactants (values found at optimum conditions) Surfactant
cmc (mM)
Gas hold-up
Half life (s)
CTAB (c)
0.9 (in H2O)
0.70 <500 40–260 (Fuda et al., 2005) (Fuda et al., 2005; (Chaphalkar Save and Pangarkar, et al., 1993) 1994; Yan et al., 2005)
Tween 20 (n-i) 0.056
0.4–0.63 (Dermiki, 2009)
100–600 (Dermiki, 2009)
Tween 60 (n-i) 0.027
0.10–0.60 (Dermiki et al., 2009)
100–500 (Dermiki et al., 2009)
Bubble size (mm)
AOT (a)
0.84 0.02–0.65 (Fuda, 2004) (Jauregi, 1997)
30–930 46–101 (Jauregi et al., 1997) (Jauregi et al., 2000)
SDS (a)
8.08
<550 (Yan et al., 2005)
0.70 (Matsuhita et al., 1992)
10.2.3 Generation of CGA Formation of aphrons The formation of CGA requires, as described by Sebba (1985), a horizontal disc that rotates at very high speeds. Baffles are also recommended in order to achieve the required mixing regime and produce smaller bubbles (Jauregi et al., 1997). Furthermore, Jauregi et al. (1997) determined the power input required to generate CGA and found that there is a minimum power requirement to generate CGA, which is dependent on the surfactant concentration as this affects the volume of air incorporated; for surfactant concentrations below and around the cmc the minimum power requirement to generate CGA was 45 kW m–3. In our laboratory, CGA are typically generated by stirring a surfactant solution with a laboratory mixer at 8000 rpm. A four-bladed impeller with a high shear screen (Fig. 10.2) was found to generate CGA dispersions with small bubble sizes (average diameters ranged between 35 and 70 mm) (Jauregi et al., 1997). Some other apparatus reported by Sebba (1985) involved the use of a venturi throat at which gas is admitted but this is not satisfactory for production of CGA on a large scale. In a recent study by Xu et al. (2008), the use of sonication for the formation of CGA was investigated and compared with mechanical agitation, which is the most widely used method. They found that sonication led to CGA with higher gas hold-up, smaller bubble size, higher number of bubbles and larger interfacial area than mechanical agitation; processing time was also shorter. However, sonication is an expensive method that cannot be easily scaled up.
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292 Separation, extraction and concentration processes
Fig. 10.2 Impeller used with laboratory mixer for the generation of CGA.
Type of surfactants used in the generation of CGA A wide range of surfactants can be used for the generation of CGA: anionic, e.g. sodium dodecyl sulfate (SDS); cationic, e.g. cetyltrimethylammonium bromide (CTAB); nonionic, e.g. Tween 20. Characteristics of CGA generated with these surfactants are summarised in Table 10.2; also an overview of the characteristics of CGA generated from different surfactant solutions is given by Jauregi et al. (1997). The use of surfactants for food applications can be problematic particularly if they are as toxic as the ionic surfactants. Nonionic surfactants on the other hand can be advantageous to use as they are non-toxic and they are used in the formulation of medicines; hence, they may not need to be removed from the CGA phase and at the same time could help to formulate the final product (section 10.3.1). The use of natural surfactants such as saponin extracted from plants has also been explored (Kommalapati et al., 1998). The biosurfactant rhamnolipid produced by Pseudomonas aeroginosa (Feng et al., 2009) is an alternative to the synthetic ones because biosurfactants are biodegradable and hence have a less harmful impact on the environment. However, these surfactants, although harmless to the environment, can, in high concentrations, be toxic so their use in food applications needs to be further studied. Moreover, not all surfactants are suitable for generation of CGA. For instance, those surfactants with poor
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Separation of value-added bioproducts by colloidal gas aphrons 293 foaming ability cannot produce CGA. Surfactants that have a tendency to form very stable micelles or vesicles have poor foaming ability. Micellar stability is inversely related to foaming ability because very stable micelles are less capable of providing the flux of surfactants necessary to stabilise the new air–solution interface created during foaming (Rosen, 2004). In experiments carried out in our laboratory with the less toxic cationic surfactant dioctadecyldimethylammonium bromide (DODAB) at 0.1 and 1 mM aqueous solutions, we could not generate CGA. This surfactant has the ability to form very stable vesicles owing to its structure and, hence, has very low foaming ability. However, it would be very useful to generate CGA using DODAB because it is used in medicinal applications and may be suitable for food applications. Span surfactants have a low foaming ability whereas their nonionic ethoxylated derivatives, Tween surfactants, have a relatively high foaming ability. CGA formed in combination with those two surfactants led to decreased CGA stability compared with the ones produced only from the Tween surfactants (Dermiki et al., 2009) (section 10.3.2, Table 10.4). Worden and Scranton (2000) describe an interesting method to produce reversible gas or liquid aphrons using a combination of polymers as surfactants such as methacrylic acid and polyethylene glycol. The method is based on the polymers changing their foaming and emulsifying properties with pH thus a fast coalescence of CGA is achieved by changing the pH. This can be of particular interest for applications such as waste-treatment and for increasing mass transfer in bioreactors. However, for bioseparations the removal of surfactant still remains a problem.
10.3 Applications of CGA in the recovery of value-added food products As a consequence of the above properties, researchers have considered various applications for CGA, with a particular focus on separation processes (Table 10.3). For example, CGA have been used for the flotation of biological products such as microbial cells and proteins; see also review by Jauregi and Varley (1999) on applications of CGA in biotechnology. Other interesting biotechnological applications include the use of CGA in bioprocesses to enhance gas mass transfer in fermentations (Bredwell and Worden, 1998; Kaster et al., 1990; Park et al., 2009) or even in wastewater treatment (Dai et al., 2004). A number of recent research studies focus as well on the environmental applications of CGA such as the remediation of soil (Boonamnuayvitaya et al., 2009; Chu, 2003; Couto et al., 2009; Kommalapati et al., 1998; Roy et al., 1995) and the removal of oil and dyes from wastestreams (Roy et al., 1992) as summarised in Table 10.3. High recoveries of proteins and other bioproducts indicate the potential of CGA in downstream processing, where alternative methods of reducing
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Table 10.3 Summary of applications of CGA in separation processes (adapted from Dermiki et al., 2008) Surfactant
Principles/driving forces of the separation
References
Recovery of antioxidants: astaxanthin
CTAB (c)
Recovery is driven by electrostatic interactions
(Dermiki et al., 2008) (Dermiki et al., 2009) (Alves et al., 2006) (Spigno and Jauregi, 2005)
Highest recovery and enrichment factor under conditions that favour electrostatic interactions. Important effect of flocculating agent
(Zidehsaraei et al., 2009) (O’Connell and Varley, 2001)
Recovery is driven by electrostatic interactions Electrostatic interactions are important for the recovery, the selectivity can be manipulated by changing conformation of proteins Continuous separation more efficient than batch Hydrophobic interactions play an important role
(Fuda et al., 2005) (Fuda et al., 2004) (Fuda and Jauregi, 2006)
Better recovery with the ionic surfactants because they produce more stable dispersions
(Mansur et al., 2006)
Recovery is driven by mechanism of bubble flotation
(Mansur et al., 2004)
The adsorption of yeast on CGA follows the Langmuir model. Changes in pH and feed concentration lead to a change of mechanism of bubble attachment and detachment from a monolayer to a multilayer adsorption. The volume of air incorporated in the system is a significant factor for separation
(Hashim et al., 1998) (Hashim et al., 1995) (Hashim et al., 2000)
norbixin gallic acid Recovery of enzymes: glucoamylase lipase immobilisation
CTAB (c) CTAB (c) TTAB (c) SDS (a)
Recovery of proteins: CTAB (c) b-lactoglobulin lactoferrin, lactoperoxidase AOT (a) CTAB (c), AOT (a) whey proteins whey proteins lysozyme, b-casein Removal of fine particles
Clarification of yeast cells
SLS (a) TWEEN 20, TWEEN 40, TWEEN 60, TWEEN 80 (n-i) HTAB (c) SDBS (a) TWEEN 20 (n-i) HTAB (c) SDBS (a) TWEEN 20 (n-i) BDHA (c)
(Amiri and Valsaraj, 2004) (Jarudilokkul et al., 2004)
294 Separation, extraction and concentration processes
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Application
Clarification of suspensions: oil mill effluent suspension of microalgae suspension of inorganic minerals
AOT (a), SDBS (c), LUX flakes (a), BDHA (c)
Recovery is enhanced when electrostatic interactions are (Basu and Malpani, 2001) favoured (Roy et al., 1992) The height of the separation column is an important factor in (Hashim et al., 1999) order to ensure the required separation CGA better oil removal efficiencies. Removal of oil positively affected by initial oil concentration and particle diameter CGA more efficient than conventional surfactant solution CGA more efficient than conventional surfactant solutions Nonionic surfactant better for the removal of naphthalene because it has better solubilising power for the compound CGA more efficient than simple waterflood, the recovery results from solubilisation of the hydrophobic compound in the surfactant Maximum effectiveness of separation was achieved at pH close to the pK of each surfactant, removal effectiveness increased when CGA were used at the temperature at which they were produced, rather that at high temperatures, the effectiveness of solids removal was a function of air-tosolids ratio and of the solids concentration
(Couto et al., 2009) (Boonammuayvitaya et al., 2009) (Roy et al., 1994) (Roy et al., 1995) (Kommalapati et al., 1998) (Subramaniam et al., 1990)
SDS, sodium dodecyl sulfate; SDBS, sodium dodecylbenzenesulfonate; SLS, sodium lauryl sulfate; AOT, sodium bis (2-ethyl hexyl) sulfosuccinate; HTAB (= CTAB), hexadecyltrimethylammonium bromide; TTAB, tetradecyltrimethylammonium bromide; BDHA, benzyldimethylhexadecylammonium chloride.
Separation of value-added bioproducts by colloidal gas aphrons 295
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Remediation of wastewater: removal of organic dyes HTAB (c), SDS (a), SDBS (a) clarification of oily SDS (a) wastewater Remediation of soil: remediation of sandy soil SDS (a) contaminated with diesel oil BioNonex and Biosolve removal of pyrene from soil SDS (a) hazardous oily waste Tergitol (n-i), HTAB flushing naphthalene (c), SDS (a) Plant-based surfactant removal of hexachlorobenzene (HCB)
296 Separation, extraction and concentration processes processing costs are continuously sought. This is particularly important in the extraction of value-added products from food and agricultural waste. Research carried out by our group shows that CGA can successfully recover proteins from whey (Fuda et al., 2005; 2004), polyphenols from wine waste extracts (Spigno and Jauregi, 2005; Spigno et al., 2010) and carotenoids such as norbixin from plant extracts (Alves et al., 2006). We have also obtained high recovery of astaxanthin particles from a fermentation mixture containing yeast cells (Dermiki et al., 2008). A good understanding of the mechanism of separation is required in order to predict and optimise separations. Therefore, we investigated the mechanism of separation of proteins with CGA generated from ionic surfactants and the mechanism of separation (flotation) of astaxanthin particles, both of which are described in the following subsections. 10.3.1 Recovery of soluble compounds and mechanism of separation Recovery of proteins Similarly to foam fractionation, CGA can be applied for the recovery of soluble compounds. For foam fractionation, the selectivity of separation relies on differences in surface tension of the mixture components whereas in CGA fractionation the selectivity is based on differences in interaction between the surfactant in CGA and the mixture components. Some studies have described the successful separation of proteins with CGA for protein separation. Jauregi and Varley (1999) and Noble and Varley (1999) detailed laboratory scale recovery of proteins from single-protein model solutions (Fig. 10.3). A few studies (Amiri and Valsaraj, 2004; Fernandes et al., 2002b; Fuda et al., 2005) reported the recovery of proteins from a crude extract partly because of a lack of understanding of the mechanism of separation. Our investigation of the separation of proteins from whey has led to an improved insight into the mechanism of separation and to identification of the main operating parameters. Whey contains many proteins that differ in physicochemical properties such as charge, size and hydrophobicity. In a study with CGA generated with an anionic surfactant AOT and whey (Fuda et al., 2004), we found that selective separation of lactoferrin and lactoperoxidase could be achieved at conditions which promote strong electrostatic interactions between the proteins and surfactant (pH= 4). Interestingly, high ionic strength led to higher purity as, at these conditions, drainage of CGA is favoured, resulting in a higher number of contaminant proteins partitioning into the liquid phase. Statistical analysis of the data led to predictive models that correlated purity and enrichment ratio (ratio of protein concentration in the aphron phase to liquid phase) with the main operating parameters which were pH, ionic strength and volume of whey; these models could explain 70% and 62% of the data for purity and enrichment ratios, respectively. Volume of whey is a measure of total protein load and it was found that an increase in volume of whey resulted
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Separation of value-added bioproducts by colloidal gas aphrons 297 Feed suspension
Aphron phase
Liquid phase CGA generator
Fig. 10.3 Schematic representation of the batch small-scale procedure for protein separation using CGA.
in lower separation. This could be partly explained in terms of capacity of CGA but also in terms of protein competition as, at high protein load, an increase in protein–protein interaction can reduce selectivity of separation. A study of separation with model mixtures of lactoferrin–lactoperoxidase (Lf–Lp) and b-lactoglobulin (b-Lg) (Fuda and Jauregi, 2006) demonstrated that the recovery of Lf–Lp decreases from 80 to 45% when in the presence of b-Lg. Further investigations into the mechanism of the separation with CTAB generated CGA and whey confirmed that CGA when generated with ionic surfactants can act similarly to ion exchangers and hence selectivity can be manipulated by the type of surfactant (cationic or anionic), pH and ionic strength of the solution. Thus, under optimum conditions and with CGA, generated by CTAB, selective separation of b-Lg from whey was obtained with 90% b-Lg recovered in the CGA phase and most of the other proteins including BSA and a-lactalbumin remained in solution at basic pH (Fuda et al., 2005). Under these conditions b-Lg interacted strongly with CTAB mainly by electrostatic interactions which resulted in its precipitation. This maximised drainage, which led to the recovery of b-Lg in the form of an insoluble complex with the surfactant and most of the other proteins in the drained liquid. Although results could be explained based on CGA acting as ion exchangers, the poor recovery of proteins with similar surface charge characteristics as b-Lg led to further investigations. Protein–surfactant interactions were further investigated by measurements of zeta potential and fluorescence and these revealed that conformational features of the protein and particularly in relation to denaturation by surfactant molecules also had an influence on the selectivity of the separation. Some studies have been carried out using nonionic surfactants for the © Woodhead Publishing Limited, 2010
298 Separation, extraction and concentration processes recovery of proteins. Jarudilokkul et al. (2004) showed that hydrophobic interactions are important when nonionic surfactants are used for the recovery of proteins such as, lysozyme and b-casein. Noble et al. (1998) applied CGA generated with Triton X-100 to a range of proteins and found that the highest recoveries (74%) were obtained with the most hydrophobic protein (thaumatin). Similarly, Fernandes et al. (2002a) found that the enzyme with a hydrophobic fusion tag was recovered with a higher yield than the wild type cutinase with CGA generated with the nonionic detergent Triton X-114, thus suggesting that hydrophobic interactions were important in the recovery of these proteins. Foam fractionation has also been applied to the recovery of proteins. In this method, bubbles are generated by sparging air through a sintered glass sparger of a given pore size and the proteins are separated based on differences in their surface activity; those with the highest surface activity adsorb to the air-liquid interface of the rising bubbles and separate from the bulk. Noble and co-workers (1998) recovered 86% casein and 25% lysozyme from a binary mixture. Noel et al. (2002) applied a semi-batch foaming process for the recovery of lactoferrin from milk and obtained 40% recovery and a separation ratio of 17.8. Foam fractionation can be a cost effective way of fractionating bioproducts as no additional chemicals are required, but selectivity depends on differences in foaming properties of the components in the mixture and these cannot be easily manipulated by operating parameters such as pH, as foaming of proteins is mainly dependent on their structural properties. Recovery of bioactive components from plant extracts Plant extracts are an important source of value-added products such as polyphenols. In addition, there is much interest in minimising waste and extracting products from agricultural and food waste. We have investigated the recovery of norbixin with CGA from an alkali extract of annatto seeds (Alves et al., 2006). The carotenoid bixin represents more than 80% of the total carotenoids in annatto seeds and is liposoluble. When extracted in alkali conditions the ester group is hydrolysed and becomes the hydrosoluble form called norbixin. It was found that optimum recoveries were achieved at conditions that favour electrostatic interactions using the cationic surfactant CTAB (94%), whereas low recovery (40%) was obtained with the anionic surfactant AOT. The molar ratio of surfactant to norbixin was found to be an important operating parameter. Interestingly maximum recoveries were obtained at CTAB-norbixin molar ratios of 3–4, whereas at higher molar ratios recovery decreased. It was postulated that CTAB and potassium norbixinate interact electrostatically to form an insoluble complex which leads to its effective separation into the CGA phase. In this process, the removal of surfactant was investigated by acidifying the recovered solution to break up electrostatic interactions between surfactant and norbixin molecules and thus reduce the solubility of norbixin which resulted in its precipitation. Bioactive compounds are also found in waste streams of food processes
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Separation of value-added bioproducts by colloidal gas aphrons 299 such as in winemaking. One example is the polyphenols present in grape marc which is a by-product of wine production. Spigno and Jauregi (2005) have investigated the recovery of gallic acid from its standard suspensions with CGA generated from a cationic surfactant. Gallic acid is a polyphenol which becomes ionised and hence negatively charged at pH > 3. Electrostatic interactions were the main driving force for the recovery as optimum recoveries were achieved with the cationic surfactant CTAB at pH = 6. It was found that this pH was also optimum for the recovery of antioxidant activity because at a basic pH loss of antioxidant activity was observed. Further research into the application of CGA to real extracts and using a flotation column showed that at optimum conditions found for the standard gallic acid, high recoveries of total polyphenols were achieved (Spigno et al., 2010). Selectivity of the separation was low as both anthocyanins (expressed as malic acid equivalents) and polyphenols (expressed as gallic acid equivalents) were equally separated into the CGA phase. This led to the use of nonionic surfactants and resulted in high recoveries of all polyphenols (unpublished results). This is an interesting outcome as the use of nonionic surfactant enables integration of the recovery and formulation steps because these surfactants are used in food formulations, e.g. food emulsions (a list of food-grade surfactants is given by Monsalve-Gonzalez and Ochomogo (2009) and a more exhaustive list is given on the website on food additives listed in section 10.6). 10.3.2 Recovery of particles and mechanism of separation Recovery of particles using CGA CGA have been applied for the recovery of fine particles as seen in Table 10.3 and, as observed in those studies, electrostatic interactions had an important effect in the recovery (Waters et al., 2008) as found for soluble compounds (10.3.1). In addition, bubble–particle interactions which are affected by the particle and bubble sizes play a significant role in the recovery of particles. The type of surfactant is important not only because of the surface charge but also because it determines the bubble size of the CGA. Moreover, the concentration of the surfactant will also determine the bubble size and the surfactant molecules available for the recovery of the particles (Mansur et al. 2006). Furthermore, parameters that affect the stability of the CGA dispersion will affect the selectivity of the separation owing to changes in drainage (Waters et al., 2008). Therefore, one can say that the removal of particles is affected by a range of parameters, and this process becomes more challenging for complex systems where more than one type of particle is present. This is illustrated in the case study of the recovery of astaxanthin. Recovery of astaxanthin and mechanism of separation There is increasing interest in the xanthophyll astaxanthin owing to its antioxidant activity and its potential benefits to human health (Hussein et al.,
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300 Separation, extraction and concentration processes 2006). Specifically, natural astaxanthin is produced mainly by two microorganisms: the yeast Phaffia rhodozyma and the microalgae Haematococcus pluvialis. We have investigated the application of CGA following an integrated approach for the recovery of astaxanthin from suspensions of cells, more specifically from the yeast Phaffia rhodozyma (Dermiki et al., 2008). The mechanism of separation was elucidated using a range of surfactants for the generation of CGA and these were subsequently applied for the recovery of astaxanthin from standard suspensions and suspensions of cells (raw extract) as shown in Table 10.4. A minimum recovery of 40% was achieved under all conditions studied, indicating that the recovery was the result of flotation of the hydrophobic astaxanthin particles. However, when electrostatic interactions were promoted at basic pH conditions, recoveries increased further up to 80%, which indicated that CGA were even more efficient in floating the negatively charged particles than the non-charged hydrophobic particles. Interestingly, when the optimum conditions for the standard suspensions of synthetic astaxanthin were applied to the recovery of natural astaxanthin from the cells of Phaffia rhodozyma, results confirmed further that electrostatic interactions were the driving force for the recovery. To elaborate, recoveries up to 97% were obtained with the cationic surfactant CTAB under basic conditions, specifically when the cells of Phaffia rhodozyma were pretreated with NaOH. The use of NaOH had a double effect: it resulted in the higher release of astaxanthin from the cells and it facilitated its subsequent recovery with CTAB. These experiments revealed that the main operating parameters of the separation were the volumetric ratio of CGA to feed or mass ratio of surfactant to total solids in the feed and the initial concentration of total solids in the feed. Specifically, recoveries increased with increased mass ratio and with decreased total solids in the feed. Scale-up with a flotation column The optimum conditions achieved on a small scale (batch), where separation took place in a beaker (the same as for the small scale recovery of proteins, see Fig. 10.3), were applied for the recovery of astaxanthin at larger scale using a flotation column set-up at two different modes (continuous and batch). The set-up for the large scale for the two different modes studied can be seen in Fig. 10.4. Flotation is a separation method that can be used on a large scale but column flotation results in general in lower yields due to the large bubble sizes generated in these columns (Nguyen and Schulze, 2004) and it is more effective for low concentrations of feed and relatively large particle size (mm range) owing to the relative non-turbulent nature of the contact of bubbles and particles (Tasdemir et al., 2007). However, high concentrations of surfactants are necessary and their high cost can restrict their application to the smaller scale (Hashim et al., 1998). CGA, as has already been highlighted above (Table 10.2) have high interfacial area owing to their small size and
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Table 10.4 Summary of the surfactants used and the main outcomes on the recovery of astaxanthin from standard suspensions and suspensions of cells
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cmc mM HLB† Max R % (standard suspensions) Optimum conditions
CTAB (cationic)
SDS (anionic)
Tween 60 (nonionic)
Tween 20 (nonionic)
Span 80 (nonionic)
0.9 – 80%
12 – 55%
0.027 14.9 75%
0.056 16.7 59%
– 4.3 50%
pH = 2 i.s. = 0.055M
Hydrophobic interactions: 20% ethanol 20% Hydrophobic interactions: 20% ethanol
Hydrophobic interactions
Tween 60/Span 80 HLB = 14.024 or HLB = 11.487 – –
Electrostatic interactions: pH = 11 Max R % (raw extract) 97% Optimum conditions Electrostatic interactions: pretreatment with NaOH; 0.2M+homogenisation; clarified Advantages High R% for both natural and synthetic astaxanthin Disadvantages
Toxic Rat: LD50 = 410 mg kg–1
– –
Less toxic than CTAB, Food grade, rat: LD50 = 1200 mg improved biovailability, kg–1 preformulation Low R% Low R% for natural astaxanthin
25% Hydrophobic interactions: Mixture with CTAB nCTAB/nTween2 = 0.1 In H2O + homgenisation Food grade, improved biovailability, preformulation Low R%
Food grade
CGA low stability, low R%
R%, recovery percentage; i.s., ionic strength; LD50, the dose required to kill half the members of a tested population. *CTAB, hexadecyltrimethylammonium bromide; SDS, sodium dodecyl sulfate; Tween 60, polyoxyethylene sorbitan monostearate; Tween 20, polyoxyethylene sorbitan monolaurate; Span 80, sorbitan mono-oleate. † HLB, hydrophile lipophile balance.
Separation of value-added bioproducts by colloidal gas aphrons 301
Surfactant type*
302 Separation, extraction and concentration processes
Flotation column Sample collection CGA feed
CGA generator (a) Feed P2
H4 H3
Flotation column CGA feed
Sample collection
H2 H1
Effluent CGA generator
P1 (b)
P3
Fig. 10.4 Schematic representation of the scale-up of astaxanthin separation with CGA using a flotation column: (a) batch operation, (b) continuous operation.
this makes them more efficient in the adsorption of particles and molecules. Moreover, because of their rheological properties (Table 10.1) they can be easily pumped and they provide increased froth stability which is desirable for flotation processes. Owing to these properties CGA are used for the recovery of a variety of products on a large scale in batch and continuous mode using a flotation column set-up. As shown in Table 10.5, they have been used for the recovery of yeast cells (Hashim et al., 1998; 1995), particles (Mansur et al., 2006; 2004; Waters et al., 2008), oil removal (Watcharasing et al., 2008) and organic dyes from water streams (Pandit and Basu, 2002; Roy et al., 1992); in these studies a range of operating parameters were investigated in order to determine the optimum conditions for the separation. For astaxanthin, the main operating parameter was the volumetric ratio of CGA to feed, which had the same effect in both continuous and batch © Woodhead Publishing Limited, 2010
Table 10.5 Review of applications of CGA in flotation column set-up (c, cationic; a, anionic; n-i, nonionic) (adapted from Dermiki et al., 2010) Surfactant
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Separation of CTAB (c) astaxanthin from cells of Phaffia rhodozyma Diesel oil removal from water
Parameters investigated Operation/ dimensions of column Csurf, FCGA, tdrainage, [TS]init, VCGA/Vast, height of the column
Branched alcohol Csurf, tstir, rpm, CNaCl; propoxylate sulfate feed with or without CGA sodium salt (a)
Removal of fine SDS (a) particles
Comparison with conventional flotation, addition of NaCl, pH, particle size
References
Continuous, batch d = 4.5 cm, h = 50 cm
(Dermiki, 2009)
Batch CGA increased froth formation and d = 5 cm, h = 120 stability, enhanced the removal and cm enrichment ratio of oil. Highest oil removal 97%
(Watcharasing et al., 2008)
Batch d = 7 cm, h = 30 cm
SDBS (a), CTAB Surfactant type, Csurf, (c), Tween 20 (n-i) particle size
Batch d = 5 cm h = 100 cm
SDBS (a), CTAB (c)
Batch d = 5 cm, h = 100 cm
Surfactant type, Csurf, FCGA, particle concentration
Main findings
Higher recovery of CuO with CGA (76.5%) than with conventional flotation (58.3%) Addition of NaCl decreases the recovery which implies that electrostatic interactions are important Increase of Csurf up to cmc increases removal efficiency, effect of particle size, no effect of charge of the particles Ionic surfactants: stable dispersions, increased removal efficiency Removal efficiency increases with FCGA and decreases with feed concentration. Better performance of the anionic surfactant due to smaller bubble size of the CGA suspension
(Waters et al., 2008)
(Mansur et al., 2006)
(Mansur et al., 2004)
Separation of value-added bioproducts by colloidal gas aphrons 303
Application
© Woodhead Publishing Limited, 2010
Application
Surfactant
Parameters investigated Operation/ dimensions of column
Main findings
References
Separation of organic dyes
SDBS (a), CTAB (c)
Surfactant type, Csurf, FCGA, CGA diameter and gas hold-up, pH, tres, Csalt
Batch d = 4 cm, h = 60 cm
(Basu and Malpani, 2001)
Surfactant type, FCGA
Batch d @ 8 cm, h @ 100 cm
Removal increases with tres and gas hold-up of CGA. No effects of Csurf above the cmc or Csalt. Higher removal at pH conditions that favour electrostatic interactions Recovery owing to electrostatic interactions, the FCGA affects the removal of dyes
Clarification of palm oil
SDS
Sparging rate of CGA, height of column (tres)
Continuous/ Separation efficiency increases with counter current increasing sparging rate and decreasing d = 5 cm, h = 100 height of the column cm
Treatment of contaminated soil
SDS (a), HTAB (c) Tergitol (n-i) Sapindus mukorossi fruit pericarp
Residence time, type of surfactant, Csurf; CGA and conventional surfactant solution, Csurf, alternating flushing media
d h d h
SDS (a)
Flushing mode, Csurf
d = 5.75 cm h = 10 cm
= = = =
5.75 cm 10 cm 5.75 cm 10 cm
Nonionic more effective. Removal of naphthalene increases with Csurf CGA more effective for removal of hexachlorobenzene (HCB). Recovery increases with Csurf, 24 to 84 times higher than water flushing No effect of alternating media on pressure build-up and removal of HCB CGA more effective. 56% removal of oily waste, 50% removal of transmission fluid CGA more effective in the down flow operation, no effect of Csurf
(Roy et al., 1992) (Hashim et al., 1999)
(Roy et al., 1995) (Kommalapati et al., 1998)
(Roy et al., 1994)
304 Separation, extraction and concentration processes
Table 10.5 Continued
BDHA (c )
Cfeed, FCGA
Continuous/ counter current d = 5 cm h = 58 cm
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pH, Cfeed, FCGA Cfeed, FCGA, operating height
Clarification of suspensions: oil mill effluent, suspension of microalgae, suspension of inorganic minerals
SDS, SDBS (a) CTAC, DTAC, CPC (c)
Surfactant type, CGA height, dispersed phase velocity, cell loading
AOT (a) SDBS (a) BDHA (c) LUX flakes (n-i)
Surfactant type, sparging rate, superficial hydraulic loading rate, air-to-solid ratio, pH, T, flotation cell design
Particle-bubble attachment in yeast flotation follows the Langmuir adsorption mode Separation efficiency = 95%, effect of air flow rate and Cfeed Separation efficiency = 90%, effect of pH at low FCGA
(Hashim et al., 2000) (Hashim et al., 1998) (Hashim et al., 1995)
Effect of FCGA stronger at low Cfeed Concentration gradient along the height of the column, at high Cfeed Continuous/ Better performance of the cationic surfactant (Save and counter current For relatively high dispersed phase Pangarkar d = 3 cm, h = 100 velocity, pH and cell loading have no effect 1995) cm Batch, d = 20 cm V = 5 dm3 Continuous, horizontal flow Continuous/ counter current vertical flow column d = 5 cm, h = 10 cm
Maximum effectiveness of separation at pH (Subramaniam et al., 1990) close to the pK of each surfactant Removal effectiveness increased when CGA were used at the temperature at which they were produced, rather than at high temperatures Effectiveness of solids removal was a function of air-to-solids ratio and of the solids concentration R = 90–94%
Separation of value-added bioproducts by colloidal gas aphrons 305
Yeast cells
306 Separation, extraction and concentration processes experiments. Higher volumetric ratios or mass ratios of surfactant to astaxanthin in the feed resulted in higher amounts of surfactant available to interact with the particles in the feed, leading to increased recovery; the optimum volumetric ratio was 12 where recovery reached 98% and the separation ratio (the astaxanthin concentration ratio in the aphron phase and in the liquid phase) was 10. An increase in the concentration of solids in the feed led to a lower separation efficiency (Dermiki et al., 2010), which is in agreement with other studies on recovery of yeast cells (Hashim et al., 1998) and particles. This is because of saturation of the aphron phase with particles. For the batch experiments, flow rate of CGA had a significant effect on the selectivity of the process. As the flow rate increased from 70 to 150 ml min–1 selectivity reached a maximum and further increase resulted in reduced selectivity. This is because the flow rate caused changes in contact time and CGA stability which had both, a positive and a negative effect on selectivity. Increased flow rate led to increased turbulence and decreased drainage. Stability must be high enough to ensure attachment of astaxanthin but not too high so that the liquid within the bubbles is drained and cell particles partition to the bulk liquid phase (Dermiki et al., 2010). In continuous operation, flow rate and height or the residence time of CGA in the column were the main operating parameters. Hence, separation increased with increasing flow rate as for the batch operation mode. For a specific flow rate, separation increased at higher levels of the column as this corresponded to longer residence times of CGA. However, the optimum flow rate for separation did not lead to the highest selectivity, which was also in agreement with the results in the batch experiments. The experiments with astaxanthin showed that because the astaxanthin particles are significantly smaller than the cells, a different mechanism is responsible for their separation in a flotation column: adsorption of small particles to the bubbles is mainly driven by electrostatic interactions, whereas attachment of the cell/cell aggregates to the bubbles occurs mainly because of non-specific interfacial forces such as buoyancy forces (Klimpel, 1998). The studies described showed that CGA produced from low surfactant concentrations can be used for the recovery of bioproducts and for astaxanthin recoveries up to 98% could be achieved using CGA generated from CTAB solutions of 0.8 mM, which is below the cmc of CTAB (Dermiki et al., 2010) Optimum recoveries of astaxanthin were obtained with the cationic surfactant CTAB, which at high concentrations could be toxic. However, its removal was possible by ultrafiltration with membranes of regenerated cellulose with 50 kDa molecular weight cut off, at low pH where electrostatic interactions between astaxanthin and CTAB were not promoted (Dermiki, 2009). Furthermore, by using diafiltration under the same pH conditions as before, it was possible to obtain CTAB concentrations in the retentate that were below the toxic limit. Thus, the final product was not considered toxic and the recovered surfactant in the permeate could be recycled and reused.
© Woodhead Publishing Limited, 2010
Separation of value-added bioproducts by colloidal gas aphrons 307 For astaxanthin the use of nonionic surfactants was investigated because they offer some important advantages. The surfactants that were investigated here are food-grade surfactants therefore their removal is not necessary at the final stage of the process. Moreover, these surfactants enhance the bioavailability of hydrophobic compounds such as astaxanthin. It must be stressed that astaxanthin is already available in emulsion form; therefore the use of nonionic surfactants can integrate the recovery and formulation steps. As shown in Table 10.4, high recoveries were obtained for synthetic astaxanthin particles in the standard suspension experiments. However, these high recoveries were not repeated for astaxanthin from the cells of Phaffia rhodozyma. These differences in recovery can be explained on the basis of the differences in size between the synthetic astaxanthin (>1 mm) and natural astaxanthin (<0.2 mm) and on the differences in stability between CGA generated from nonionic and ionic surfactants (Dermiki, 2009). Nonionic surfactants form less stable CGA dispersions that are not appropriate for the recovery of small particles such as the natural astaxanthin particles. Despite the fact that nonionic surfactants used on their own showed relatively low recoveries of natural astaxanthin particles, recovery and selectivity increased when the nonionic surfactants were combined with the cationic CTAB (Dermiki, 2009). The use of nonionic surfactants can be advantageous and should be investigated further for the recovery of astaxanthin from other sources, such as the microalgae Haematococcus pluviallis. In this instance, astaxanthin is esterified therefore hydrophobic interactions with the nonionic surfactant are important and promoting these could lead to high recoveries.
10.4 Feasibility of industrial application of CGA The case studies presented in this chapter show that CGA can be applied to the separation of a range of value-added products from waste streams, as well as from fermentation broths. In this section the feasibility of CGA for these industrial applications is explored. An economical evaluation of CGA separation and other conventional separations used in the recovery of astaxanthin at industrial scale was carried out. The economical evaluation of CGA was compared with that of solvent extraction (SE) and supercritical carbon dioxide extraction (SCDE) in terms of operating cost and cost of equipment (Dermiki, 2009). Process economics were evaluated using the software SuperPro Designer (Intelligen, US). One of the main outcomes of this study was that when comparing the purchase cost of equipment, CGA was the most expensive. This can be attributed to the high cost of diafiltration for the removal of surfactant; if this step was removed, the equipment cost was about the same as that of supercritical carbon dioxide extraction and lower than that of solvent extraction. However, another interesting outcome was that the latter had the © Woodhead Publishing Limited, 2010
308 Separation, extraction and concentration processes highest total operational cost, mainly owing to the high cost of raw materials, followed by supercritical carbon dioxide extraction. Although surfactants are also very expensive, low concentrations are required to generate CGA and surfactants cost less than solvents. Furthermore, the cost of utilities, which includes the cost of electricity, is high for CGA extraction but it is not higher than that of SCDE despite the high energy input requirement for CGA generation. In SCDE, the electricity cost for the compression of carbon dioxide increases the overall cost of utilities. It was also interesting to find that the cost of waste disposal was the highest for SE, very similar to that of SCDE and the lowest was for CGA. In summary, although this was an approximate estimation of the economics of a CGA separation and, therefore, several assumptions had to be made, it is still a very valuable evaluation as it highlights the strengths and weaknesses of CGA separations in relation to other conventional industrial separations. From this, it can be concluded that if CGA are applied following an integrated approach, i.e. if removal of surfactant is not required, this separation could be more advantageous than others as it is more cost effective. Furthermore, CGA separation is a more environmentally friendly separation process than solvent extraction and also uses less hazardous equipment than supercritical carbon dioxide extraction.
10.5 Future trends As shown above one of the drawbacks of the use of CGA to separate bioproducts is the cost of the surfactant and its toxicity. Removal of the surfactant can be achieved in some instances as shown in the astaxanthin study by membrane filtration/diafiltration. Although this allows recycling of the surfactant, it leads to an overall increase in the operating cost thus its application becomes less attractive. For some applications nonionic surfactants can be applied in place of the more toxic ionic surfactants and thus further process integration can be achieved as recovery of the product in an aqueous solution of a nonionic surfactant may result also in its preformulation. Therefore, the use of nonionic surfactants should be explored further. We investigated the use of nonionic surfactants for the recovery of astaxanthin but low recoveries were achieved (<20%). However, recovery and selectivity increased slightly when the nonionic surfactant was combined with the cationic surfactant CTAB (Dermiki, 2009). This is a way of decreasing the amount of the cationic surfactant. An interesting alternative to the toxic surfactants are the biosurfactants because these are nontoxic, biodegradable and most importantly they can be produced from several inexpensive waste sources making them an alternative for many synthetic surfactants (Singh et al., 2007). For CGA to be competitive against other separations, the integrative approach has to be possible as shown for astaxanthin. The application of CGA directly to the suspension of cells resulted in high recoveries and partial selective separation. © Woodhead Publishing Limited, 2010
Separation of value-added bioproducts by colloidal gas aphrons 309 As demonstrated also by the economic evaluation above, reduction in the number of process steps will reduce the overall cost. Similarly, Zidehsarei et al. (2009) generated CGA together with the feed suspension, which consisted of a solid fermentation suspension with 50% wet bran, Aspergillus niger and the product, glucoamylase. This method enabled them to simultaneously extract (from the solid substrate and biomass) and recover the enzyme in the drained liquid with recovery of 91% and specific activity of 10. In addition, they found that using aluminium sulfate as a flocculating agent they achieved improved separation of the solids in the CGA and improved drainage, which resulted in higher purity. So these applications show that CGA could be advantageous as a bioseparation for some of the value-added products and may be able to find industrial applications where high yield rather than purity is a requirement. Adsorption-based processes such as chromatography are more selective and therefore CGA cannot compete with these. Other interesting applications of CGA include environmental applications such as the remediation of soil contaminated with diesel oil (Couto et al., 2009), removal of organic dyes from wastewaters (Roy et al., 1992) and the removal of pyrene from soil with biodegradable surfactants (BioSolve and BioNex) (Boonamnuayvitaya et al., 2009). The latter shows that CGA can be generated with biodegradable surfactants and thus they are suitable for environmental applications such as remediation. This is also shown in the work by Park and co-workers (2009) who applied CGA made with a biodegradable surfactant saponin, to biodegradation of phenanthrene. They used CGA to enhance oxygen transfer during the fermentation process. Furthermore, other applications of CGA include their use to enhance ultrasound imaging owing to their reflection properties (Wheatley et al., 2006), the application as drug delivery agents using the biocompatible synthetic polymer poly(vinyl alcohol) (PVA) (Cavalieri et al., 2005) and their application as a template for the formation of tungsten oxide nanorods (Abdullah et al., 2006).
10.6 Sources of further information and advice Rosen M J (2004) Surfactants and interfacial phenomena, New Jersey, John Wiley and Sons., Inc. Sebba F (1987) Foams and biliquid foams–aphrons. Chichester, John Wiley & Sons Ltd. Nguyen A V and Schulze H J (2004) Colloidal science of flotation, New York, Marcel Dekker, Inc. United States Code of Federal Regulations (CFR) (2009) 21CFR172–Part 172 ‘Food additives permitted for direct addition to food’ Available from: http://frwebgate.access.gpo.gov/cgi-bin/multidb.cgi.
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310 Separation, extraction and concentration processes
10.7 References Abdullah SF, Radiman S, Abd. Hamid MA and Ibrahim NB (2006), ‘Effect of calcination temperature on the surface morphology and crystallinity of tungsten (VI) oxide nanorods prepared using colloidal gas aphrons method’, Colloids and Surfaces A: Physicochemical and Engineering Aspects, 280(1–3), 88–94. Alves RW, de Souza AAU, de Souza SMD and Jauregi P (2006) ‘Recovery of norbixin from a raw extraction solution of annatto pigments using colloidal gas aphrons (CGAs)’, Separation and Purification Technology, 48(2), 208–213. Amiri MC and Valsaraj KT (2004), ‘Effect of gas transfer on separation of whey protein with aphron flotation’, Separation and Purification Technology, 35, 161–167. Amiri MC and Woodburn ET (1990), ‘A method for the characterization of colloidal gas aphrons’, Transactions IChemE, 68, 154–160. Basu S and Malpani PR (2001), ‘Removal of methyl orange and methylene blue from water using colloidal gas aphrons – effect of process parameters’ Separation Science and Technology, 36, 2997–3013. Bhatia D, Goel G, Bhimania SK and Bhaskarwar AN (2005), ‘Characterization and drainage kinetics of colloidal gas aphrons’, AIChE Journal, 51, 3048–3058. Boonamnuayvitaya V, Jutaporn P, Sae-ung S and Jarudilokkul S (2009), ‘Removal of pyrene by colloidal gas aphrons of a biodegradable surfactant’, Separation and Purification Technology, 68, 411–416. Bredwell MD, Telcenhoff MD and Worden RM (1995), ‘Formation and coalescence properties of microbubbles’, Applied Biochemistry and Biotechnology, 51/52, 501–508. Bredwell M and Worden R (1998), ‘Mass-transfer properties of microbubbles. 1. Experimental studies’, Biotechnology Progress, 14, 31–38. Brookey TF and Tommy F (2004) ‘Aphron-containing well drilling and servicing fluids’. US patent 6716797. Cavalieri F, El Hamassi A, Chiessi E and Paradossi G (2005), ‘Stable polymeric microballoons as multifunctional device for biomedical uses: synthesis and characterization’, Langmuir, 21, 8758–8764. Chaphalkar PG, Valsaraj KT and Roy D (1993), ‘A study of the size distribution and stability of colloidal gas aphrons using a particle size analyser’, Separation Science and Technology, 28, 1287–1302. Chu W (2003), ‘Remediation of contaminated soils by surfactant-aided soil washing’, Practice Periodical of Hazardous, Toxic, and Radioactive Waste Management, 7, 19–24. Couto HJB, Massarani G, Biscaia Jr EC and Sant’Anna Jr GL (2009), ‘Remediation of sandy soils using surfactant solutions and foams’, Journal of Hazardous Materials, 164, 1325–1334. Dai Y and Deng T (2003), ‘Stabilization and characterization of colloidal gas aphron dispersions’, Journal of Colloid and Interface Science, 261(2), 360–365. Dai Y, Deng T, Wang J and Xu K (2004), ‘Enhancement of oxygen gas–liquid mass transfer with colloidal gas aphron dispersions’, Colloids and Surfaces A: Physicochemical and Engineering Aspects, 240, 165–171. Dermiki M, Gordon MH and Jauregi P (2008), ‘The use of colloidal gas aphrons as novel downstream processing for the recovery of astaxanthin from cells of Phaffia rhodozyma’, Journal of Chemical Technology and Biotechnology, 83, 174–182. Dermiki M (2009), ‘Recovery of astaxanthin using colloidal gas aphrons’, PhD thesis, Reading University, UK. Dermiki M, Bourquin A-L and Jauregi P (2010), ‘Separation of astaxanthin from cells of Phaffia rhodozyma using colloidal gas aphrons (CGA) in a flotation column’, Biotechnology Progress, 26(2), 477–487. Dermiki M, Gordon MH and Jauregi P (2009), ‘Recovery of astaxanthin using colloidal
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Separation of value-added bioproducts by colloidal gas aphrons 311 gas aphrons (CGA): a mechanistic study’, Separation and Purification Technology, 65, 54–64. Feng W, Singhal N and Swift S (2009), ‘Drainage mechanism of microbubble dispersion and factors influencing its stability’, Journal of Colloid and Interface Science, 337, 548–554. Fernandes S, Hatti-Kaul R and Mattiasson B (2002a), ‘Selective recovery of lactate dehydrogenase using affinity foam’, Biotechnology and Bioengineering, 79, 472– 480. Fernandes S, Mattiasson B and Hatti-Kaul R (2002b), ‘Recovery of recombinant cutinase using detergent foam’, Biotechnology Progress, 18, 116–123. Fuda E, Bhatia D, Pyle DL and Jauregi P (2005), ‘Selective separation of b-lactoglobulin from sweet whey using CGAs generated from a cationic surfactant CTAB’, Biotechnology and Bioengineering, 90, 532–542. Fuda E, Jauregi P and Pyle DL (2004), ‘Recovery of lactoferrin and lactoperoxidase from sweet whey using colloidal gas aphrons (CGAs) generated from an anionic surfactant, AOT’, Biotechnology Progress, 20, 514–525. Fuda E and Jauregi P (2006) ‘An insight into the mechanism of protein separation by colloidal gas aphrons (CGA) generated from ionic surfactants’, Journal of Chromatography B, 843, 317–326. Hashim MA, Dey A, Hasan S and Sen Gupta B (1999), ‘Mass transfer correlation in flotation of palm oil by colloidal gas aphrons’, Bioprocess Engineering, 21, 401–404. Hashim MA, Kumar SV and SenGupta B (2000), ‘Particle-bubble attachment in yeast flotation by colloidal gas aphrons’, Bioprocess and Biosystems Engineering, 22, 333–336. Hashim MA, Sen Gupta B, Kumar VS, Lim R, Lim SE and Lim C (1998), ‘Effect of air to solid ratio in the clarification of yeast by colloidal gas aphrons’, Journal of Chemical Technology & Biotechnology, 71, 335–339. Hashim MA, Sen Gupta B and Subramaniam MB (1995), ‘Investigations on the flotation of yeast cells by colloidal gas aphrons (CGA)’, Bioseparation, 5, 167–173. Hussein G, Sankawa U, Goto H, Matsumoto K and Watanabe H (2006), ‘Astaxanthin, a carotenoid with potential in human health and nutrition’, Journal of Natural Products, 69, 443–449. Jarudilokkul S, Rungphetcharat K and Boonamnuayvitaya V (2004), ‘Protein separation by colloidal gas aphrons using nonionic surfactant’, Separation and Purification Technology, 35, 23–29. Jauregi, P (1997), ‘Colloidal gas aphrons (CGA) a novel approach to protein recovery’, PhD thesis, Reading University, UK. Jauregi P, Gilmour S and Varley J (1997), ‘Characterisation of colloidal gas aphrons for subsequent use for protein recovery’, Chemical Engineering Journal, 65, 1–11. Jauregi P, Mitchell GR and Varley J (2000), ‘Colloidal gas aphrons (CGA): dispersion and structural features’, AIChE Journal, 46, 24–36. Jauregi P and Varley J (1999), ‘Colloidal gas aphrons: potential applications in biotechnology’, Trends in Biotechnology, 17, 389–395. Kaster JA, Michelsen DL and Velander WH (1990), ‘Increased oxygen transfer in a yeast fermentation using a microbubble dispersion’, Applied Biochemistry and Biotechnology, 24/25, 469–484. Klimpel R (1998), ‘Introduction to solid–solid separation of fine particles’, Florida, NSF Engineering Research Center for Particle Science and Technology. Kommalapati RR, Valsaraj KT, Constant WD and Roy D (1998), ‘Soil flushing using colloidal gas aphron suspensions generated from a plant-based surfactant’, Journal of Hazardous Materials, 60, 73–87. Larmignat S, Vanderpool D, Lai HK and Pilon L (2008), ‘Rheology of colloidal gas aphrons (microfoams)’, Colloids and Surfaces A: Physicochemical and Engineering Aspects, 322, 199–210.
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312 Separation, extraction and concentration processes Mansur EHA, Wang Y and Dai Y (2004), ‘Separation of fine particles by using colloidal gas aphrons’, Chinese Journal of Chemical Engineering, 12, 286–289. Mansur EHA, Wang Y and Dai Y (2006), ‘Removal of suspensions of fine particles from water by colloidal gas aphrons (CGAs)’, Separation and Purification Technology, 48, 71–77. Matsushita K, Mollah AH, Stuckey DC, del Cerro C and Bailey AI (1992), ‘Predispersed solvent extraction of dilute products using colloidal gas aphrons and colloidal liquid aphrons: aphron separation, stability and size’, Colloids and Surfaces, 69, 65–72. Monsalve-Gonzalez A and Ochomogo M (2009) ‘Natural flavour enhancement compositions for food emulsions’, US patent 0196972. Moshkelani M and Amiri MC (2008), ‘Electrical conductivity as a novel technique for characterization of colloidal gas aphrons (CGA)’, Colloids and Surfaces A: Physicochemical and Engineering Aspects, 317, 262–269. Nguyen AV and Schulze HJ (2004), ‘Industrial Applications’, in Nguyen AV and Schulze HJ, Colloidal Science of Flotation, New York, Marcel Dekker, 813–842. Noble M, Brown A, Jauregi P, Kaul A and Varley J (1998), ‘Protein recovery using gas–liquid dispersions’, Journal of Chromatography B, 711, 31–43. Noble MJ and Varley J (1999), ‘Colloidal gas aphrons generated from the anionic surfactant AOT for the separation of proteins from aqueous solution’, Journal of Chemical Technology and Biotechnology, 74(3), 231–237. Noel J, Ales P and Tanner R (2002), ‘Foam fractionation of a dilute solution of bovine lactoferrin’, Applied Biochemistry and Biotechnology, 98–100, 395–402. O’Connell PJ and Varley J (2001), ‘Immobilization of Candida rugosa lipase on colloidal gas aphrons (CGAs)’, Biotechnology and Bioengineering, 74, 264–269. Pandit P and Basu S (2002), ‘Removal of organic dyes from water by liquid–liquid extraction using reverse micelles’, Journal of Colloid and Interface Science, 245, 208–214. Park JY, Choi YJ, Moon S, Shin DY and Nam K (2009), ‘Microbubble suspension as a carrier of oxygen and acclimated bacteria for phenanthrene biodegradation’, Journal of Hazardous Materials, 163, 761–767. Rosen M (2004), ‘Foaming and antifoaming by aqueous solutions of surfactants’, Surfactants and interfacial phenomena, New Jersey, John Wiley & Son. Roy D, Kommalapati RR, Valsaraj KT and Constant WD (1995), ‘Soil flushing of residual transmission fluid: application of colloidal gas aphron suspensions and conventional surfactant solutions’, Water Research, 29, 589–595. Roy D, Valsaraj KT, Constant WD and Darji M (1994), ‘Removal of hazardous oily waste from a soil matrix using surfactants and colloidal gas aphrons under different flow conditions’, Journal of Hazardous Materials, 38, 127–144. Roy D, Valsaraj KT and Kottai SA (1992), ‘Separation of organic dyes from wastewater by using colloidal gas aphrons’, Separation Science and Technology, 27, 573–588. Save SV and Pangarkar VG (1994), ‘Characterisation of colloidal gas aphrons’, Chemical Engineering Communications, 127, 35–54. Save SV and Pangarkar VG (1995), ‘Harvesting of Saccharomyces cerevisiae using colloidal gas aphrons’, Journal of Chemical Technology and Biotechnology, 62, 192–199. Sebba F (1972), ‘Biliquid foams – a preliminary report’ Journal of Colloid and Interface Science, 40, 468–474. Sebba F (1985), ‘An improved generator for micron-sized bubbles’, Chemistry and Industry, 3, 91–92. Sebba F (1987), Foams and Biliquid Foams–Aphrons. Chichester, John Wiley & Sons Ltd. Singh A, Van Hamme JD and Ward OP (2007), ‘Surfactants in microbiology and biotechnology: Part 2. Application aspects’. Biotechnology Advances, 25, 99–121. Spigno G, Dermiki M, Pastori C, Casanova F and Jauregi P (2010), ‘Recovery of gallic acid with colloidal gas aphrons generated from a cationic surfactant’, Separation and Purification Technology, 71, 56–62. © Woodhead Publishing Limited, 2010
Separation of value-added bioproducts by colloidal gas aphrons 313 Spigno G and Jauregi P (2005), ‘Recovery of gallic acid with colloidal gas aphrons (CGA)’, International Journal of Food Engineering, 1, 1–10. Subramaniam MB, Blakebrough N and Hashim MA (1990), ‘Clarification of suspensions by colloidal gas aphrons’, Journal of Chemical Technology and Biotechnology, 48, 41–60. Tasdemir A, Tasdemir T and Öteyaka B (2007), ‘The effect of particle size and some operating parameters in the separation tank and the downcomer on the Jameson cell recovery’, Minerals Engineering, 20, 1331–1336. Tseng H, Pilon L and Warrier GR (2006), ‘Rheology and convective heat transfer of colloidal gas aphrons in horizontal mini-channels’, International Journal of Heat and Fluid Flow, 27, 298–310. Watcharasing S, Angkathunyakul P and Chavadej S (2008), ‘Diesel oil removal from water by froth flotation under low interfacial tension and colloidal gas aphron conditions’, Separation and Purification Technology, 62, 118–127. Waters KE, Hadler K and Cilliers JJ (2008), ‘The flotation of fine particles using charged microbubbles’, Minerals Engineering, 21, 918–923. Wheatley M, Forsberg F, Dube N, Patel M and Oeffinger B (2006), ‘Surfactant-stabilized contrast agent on the nanoscale for diagnostic ultrasound imaging’, Ultrasound in Medicine and Biology, 32, 83–93. Worden RM and Scranton AB (2000), ‘Method for forming reversible colloidal gas or liquid aphrons and compositions produced’, US patent 6022727. Xu Q, Nakajima M, Ichikawa S, Nakamura N and Shiina T (2008), ‘A comparative study of microbubble generation by mechanical agitation and sonication’, Innovative Food Science & Emerging Technologies, 9, 489–494. Yan Y-L, Qu C-T, Zhang N-S, Yang Z-G and Liu L (2005), ‘A study on the kinetics of liquid drainage from colloidal gas aphrons (CGAs)’, Colloids and Surfaces A: Physicochemical and Engineering Aspects, 259, 167–172. Zhao J, Pillai S and Pilon L (2009), ‘Rheology of colloidal gas aphrons (microfoams) made from different surfactants’, Colloids and Surfaces A: Physicochemical and Engineering Aspects, 348, 93–99. Zidehsaraei AZ, Moshkelani M and Amiri MC (2009), ‘An innovative simultaneous glucoamylase extraction and recovery using colloidal gas aphrons’, Separation and Purification Technology, 67, 8–13.
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(a)
(b)
(c)
Plate I Separation of methyl orange from methylene blue using CGA from the cationic surfactant CTAB. (a) Green solution: mixture of an anionic dye (methyl orange) and a cationic dye (methylene blue). (b) Mixture of the dyes with CGA generated from the cationic surfactant CTAB. (c) After the separation: methyl orange partitions to the aphron phase (yellow aphrons) and methylene blue to the liquid phase.
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314 Separation, extraction and concentration processes
11 Membrane bioreactors and the production of food ingredients M.-P. Belleville, D. Paolucci-Jeanjean and G. M. Rios, European Institute of Membranes, France
Abstract: Membrane bioreactors (MBRs) and their potential uses in food processing and food ingredients production are discussed, where ‘membrane bioreactor’ applies to any system constituted from a reactor working with enzymes or whole cells as the catalyst and from a membrane for separation and/or contacting purposes. The various types of MBR are presented and their advantages as well as their drawbacks are discussed. Several applications of MBRs in different food areas are then presented. Key words: membrane bioreactors, enzymatic membrane reactors, whole cell membrane reactors, bioproduction, food processing.
11.1 Introduction With increasing awareness of environmental and cost issues, biotransformations are gaining ground rapidly owing to the advantages that they offer over conventional technologies. Enzyme catalyzed reactions occur with high rates at room temperature (thermal degradation of labile compounds is avoided) with a minimum use of chemicals and a reduced number of reaction steps by avoiding by-products. Furthermore, owing to recent food scares, consumers’ fears of artificial products have increased the interest in biotransformation processes, which lead to products regarded as natural ingredients. Generally, bioconversions are carried out in classical batch reactors, which are easy to implement at any scale but present several disadvantages, in particular on the industrial scale: low efficiency owing to start up and shut down procedures, high labour costs and great variability of product quality owing to batch-to-batch variations. Moreover, at the end of the reaction,
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Membrane bioreactors and the production of food ingredients 315 biocatalysts are removed even if they are still active entailing high processing costs in particular when enzymes are used as catalyst. To reduce enzyme cost, enzyme immobilization on solid supports has been widely used. Thus, at the end of reaction, the catalyst can be easily separated from the reaction medium, cleaned and reused in further reactions. Moreover immobilized enzymes can be used in semicontinuous or continuous processes at the industrial stage because immobilization improves their stability (Iyer and Ananthanarayan, 2008) as well as other properties such as specificity and selectivity (Mateo et al., 2007). The operational costs are then reduced. The membrane reactor that associates reaction and membrane separation can represent an attractive alternative approach to classical methods of biocatalyst immobilization. In such a reactor, the membrane ensures the complete rejection of the enzyme or cells in order to maintain the full activity inside the reacting volume without the need for other immobilization techniques while the products are continuously extracted from the medium, thus reducing their inhibitory effects on the reaction rate. The first studies on membrane reactors applied to biological or food processes, here referred to as membrane bioreactors (MBRs), were carried out in the 1970s and since then numerous reviews have been published (Belfort, 1989; Cheryan and Mehaia, 1986; Giorno and Drioli, 2000; Prazeres and Cabral, 1994; Rios et al., 2004; Sanchez Marcano and Tsotsis, 2002). This chapter focuses on membrane bioreactors and their potential uses in food processing and food ingredients production. Firstly, the various types of MBR are presented and their advantages as well as their drawbacks are discussed. Then, several applications of MBRs in different food areas are given.
11.2 Membrane bioreactors for the production of food ingredients Membrane bioreactors correspond to the association of a reactor involving enzymes or whole cells as catalyst with a membrane separation unit. However, MBRs can be classified into two types depending on the role played by the membrane. The first type, also called free enzymes or cells membrane bioreactors (FEMBR or FCMBR) concerns MBRs involving free enzymes or cells and in that instance the function of the membrane is to retain the biocatalysts inside the reactor throughout the process. In the second type, the biocatalyst is immobilized on the membrane surface or within its pore structure and the reaction occurs at the outside or internal surface of the material during the membrane transfer.
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316 Separation, extraction and concentration processes 11.2.1 Free enzymes or free cells membrane bioreactors (FEMBRs or FCMBRs) The first configuration (Fig. 11.1) still remains the most popular and simplest one. In the continuous mode, a stirred tank reactor continuously fed with fresh substrate solution is associated with a membrane separation unit. In order to limit polarization phenomena and membrane fouling, the reaction medium generally flows tangentially along the membrane before being recycled into the reactor. Such reactors are also referred to as continuous recycle membrane bioreactors (CRMBR). The choice of the membrane is essential as regards performance; the membrane molecular weight cut off (MWCO) should be chosen in order to ensure the retention of enzymes or cells as well as the substrate; but the membrane pores must be large enough to enable the product to pass through the membrane module. As many enzymes have a molecular weight of 10 to 80 kDa, ultrafiltration membranes with a molecular cut-off between 1 and 100 kDa are generally used in FEMBR. For FCMBRs, microfiltration membranes are generally preferred. Chèze-Lange et al. (2002) reported that pore diameters should be below 1.4 mm, in order to prevent cell leakage. For the membrane material, a wide range of polymeric membranes, including polysulfone, polyethersulfone, polyamide and polypropylene, is used as well as inorganic membranes. Although mineral membranes are more expensive than organic ones, they are often more attractive because they can endure higher temperatures and pressures as well as strong chemical treatments for regeneration or even sterilization. Particular attention should be paid to the selection of the membrane material because electrostatic or hydrophobic interactions between the enzymes or cells and the membrane could affect the process performance. In particular, enzyme or cell adsorption leads to a reduction of permeation flux owing to membrane fouling. Moreover, as reported by Bódalo et al. (2004) and Paolucci-Jeanjean et al. (2001), enzyme
Substrate
Products
Membrane unit
Biocatalysts
Reactor unit
Fig. 11.1 Continuous recycle membrane bioreactor (CRMBR).
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Membrane bioreactors and the production of food ingredients 317 adsorption onto the membrane surface may induce a catalytic efficiency decrease in FEMBRs. In order to improve the control of fouling Kwon et al. (2006) suggest the use of submerged membrane reactor (Fig. 11.2). This MBR configuration, more generally called immersed membrane bioreactor (iMBR), is widely used for aerobic wastewater treatment (Judd, 2008). In practice, the membrane is placed inside the bioreactor and permeate is obtained thanks to suction force or to the transmembrane pressure owing to hydrostatic pressure. In such reactors, fouling is avoided by the use of coarse-bubble aeration which generates bubbles close to the membrane surface. iMBRs are also more efficient in regard to cells recycling and power consumption. Although 30 to 50% of the energy demand is caused by aeration of the membrane, high air fluxes promote higher permeate flux and thus lower the required membrane areas. Furthermore, the use of intermittent air flux or optimized design of the membrane module can reduce the energy consumption (Judd, 2008). In a third type of FEMBRs or FCMBRs: hollow-fibre MBRs, biocatalysts (enzymes or cells) are retained on the shell side whereas substrate flows along the lumen side of the membrane (Krastanov et al., 2007; Novalin et al., 2005). In such reactors, mass transfer resistances are high; the substrate has to be transferred across the membrane to the shell side where the biocatalyst is located and then the product has to be transferred back to the lumen side of the membrane. In addition to the reactors described above, it is also possible to couple a bioreactor to a pervaporation unit. For enzymatic ester synthesis, this coupling enables an effective in situ removal of water (a by-product of the reaction) through hydrophilic pervaporation membranes. In this way, the reaction equilibrium is displaced in favour of synthesis and the conversion rate is increased (Bartling et al., 2001; Won et al., 2006; Ziobrowski et al., 2009). In a recent review, Vane (2005) detailed the interest of such coupling for product recovery from biomass fermentation processes. The pervaporation unit can be directly associated with the reactor except for reactions involving thermosensitive cells or enzymes. In practice, the performance is improved Products Substrate
Membrane Cells Reactor
Fig. 11.2 Submerged or immersed membrane bioreactor (iMBR).
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318 Separation, extraction and concentration processes when the system operates at high temperature. The non-thermoresistant biocatalysts need to be removed before being subjected to denaturating conditions. In that instance, the pervaporation module can be fed with the permeate obtained from a CRMBR (Del Amor Villa and Wichmann, 2005). To avoid product or by-product inhibition, pervaporation can also be replaced by perstraction or electrodialysis. Qureshi and Maddox (2005) combined a batch reactor with a perstraction unit in order to produce butanol from concentrated lactose-whey. Valadez-Blanco et al. (2008) associated a bioreactor with a nanofiltration membrane contactor in order to extract continuously the R-citronellol produced by baker’s yeast from geraniol in hexane. This MBR configuration avoided the formation of emulsions, thus reducing downstream separation and enabling increased substrate loadings. Meynial-Salle et al. (2008) developed an integrated membrane–bioreactor–electrodialysis system that permits the bioproduction of succinic acid at high concentration, productivity and yield using Anaerobiospirillum succiniciproducens. Compared with classical reactor configurations such as batch, fixed or fluidized-bed reactors, free enzyme or cell MBRs offer some advantages. They permit a continuous operating mode and the high concentration of biocatalyst as well as the continuous removal of inhibitors. This entails high production rates, which ensure the economic viability of the process. In addition, as reaction and separation zones are placed in series, they can be dealt quite independently in order to optimize the performance of the whole process. Consequently, production can be optimized by acting separately on temperature, pH, substrate and biocatalyst concentrations, fluid velocity, individual control of hydraulic and biomass residence times, pressure, reactor volume or membrane surface. Unfortunately, such MBRs also present some drawbacks. The first problem relates to the change in biocatalyst activity. The loss of activity is essentially owing to adsorption onto membrane surface as well as mechanical stress, which entails enzyme and cell deactivation. However, for reactors involving cells, a part of this loss may be caused by nutrient limitation or toxin accumulation in the broth. The second main drawback is membrane fouling, which reduces the permeate flux and thus the cost-efficiency of the process. In a recent review, Meng et al. (2009) present the recent and current developments concerning the fouling behaviour, fouling factors and fouling control strategies in MBRs. Finally, if the use of FEMBR is particularly recommended when enzyme and substrate are larger than products (as in hydrolysis reactions), they are not adapted for carrying out synthesis when substrates and products show similar sizes. In this latter instance, it is better to use a reactor where the biocatalysts are immobilized on or within the membrane. Such membranes, also named ‘active membranes’, are at the heart of the so-called immobilized enzyme membrane bioreactor (IEMBR) described here below.
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Membrane bioreactors and the production of food ingredients 319 11.2.2 Immobilized enzyme membrane bioreactor (IEMBR) There are two modes for operating active membranes and associated IEMBRs as presented in Fig. 11.3. In the first mode (Fig. 11.3a), the substrate solution flows through the membrane to the biocatalyst as a result of a transmembrane pressure. The reaction occurs at the wall and the product is recovered in the permeate. The membrane is a quite specific macrosystem resulting from the assembly of swarms of microsystems (the pores), which can be regarded as microreactors. In such reactors, the mass transfer path is reduced, the contact between substrate molecules and catalysts are thus favoured. A precise control of the reaction with minimized substrate and catalyst losses, faster reactions and higher yields and cleaner products can be expected (Rios et al., 2004). In addition as the membrane thickness is much smaller than the bed length, the pressure drop and energy costs are dramatically reduced compared with fixed or fluidized bed reactors. In the second mode (Fig. 11.3b), reaction takes place at a phase-contacting interface within the membrane material. Two solutions (the substrate solution and an extracting solution that is not miscible with the substrate one) flow along each membrane side. The substrate diffuses from the feed solution, reacts with the biocatalyst and the product diffuses towards the extracting solution with which it presents a very high affinity. This biphasic reactor is particularly attractive when products and substrate show different solubilities in water and organic media as is the case in lipid hydrolysis (Knezevic et al., 2004; Merçon et al., 2000; Wang et al., 2008) or lipase-catalyzed kinetic resolution of racemic solution (Long et al., 2005; Ong et al., 2008; Wang et al., 2007). The reaction can occur without requiring emulsion formation and the product can be obtained in a single phase. However, compared with FEMBRs, the performance of such a reactor is limited by a high mass transfer resistance; indeed mass transport occurs by diffusion. Whatever the IEMBR considered, one of the important considerations is the proper incorporation of the active catalyst on or within the membrane. The three main types of active membrane preparation are shown in Fig. 11.4.
m zy
Substrate
En
Permeate
Aqueous phase
Enzyme
Feed
e
Retentate
Organic phase
Products
(a)
(b)
Fig. 11.3 Immobilized enzymatic membrane bioreactors (IEMBRs): monophasic reactor (a); biphasic reactor (b).
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320 Separation, extraction and concentration processes The biocatalyst can be (1) entrapped within the membrane structure, (2) immobilized by gelification on the membrane surface or (3) attached through covalent or noncovalent binding at the membrane surface. Entrapment within the polymeric structure can be achieved by mixing the enzyme solution with the polymeric solution used for membrane formation. The biocatalyst can be simply physically entrapped or it can be covalently linked to the polymer matrix to avoid enzyme leakage (Kanwar and Goswami, 2002; Tan et al., 2002). It is also possible to add particulate support (i.e. active carbon) previously loaded with enzymes in the polymeric solution before casting the polymeric film (Torras et al., 2008). However, the most widely used method to achieve enzyme entrapment is filtration. The enzyme solution is filtered from the support to the separating layer in order to retain the biocatalyst in the porosity of the membrane support (Sousa et al., 2001; Wang et al., 2008; Xu et al., 2006). This is a very simple procedure that leads to high enzyme loading but also has high leaching risks. Such active membranes are generally used when the feed solution permeates from the outside to the inside of the membrane or for the biphasic reactor. In this latter instance, the low solubility of proteins in organic solvent prevents their desorption. In order to limit the risk of enzyme leakage and improve the enzyme stability, Hilal et al. (2004) suggest embedding crosslinked enzyme aggregates (CLEAs) within membrane porosity. The CLEAs are formed inside the membrane pores previously filled with enzyme solution by precipitation using organic solvent with simultaneous crosslinking by glutaraldehyde. It is worth noting that active membranes elaborated by entrapment are more convenient for the reaction-limited regime rather than for a diffusion-limited regime. Active membranes can also be obtained by enzyme gelification on the membrane surface. The enzyme solution is filtered from the active side to the porous support on an ultra- or microfiltration membrane and the rejected enzymes form a gel layer on the membrane surface (Sakaki et al., 2001; Trusek-Holownia and Noworyta, 2007). The stability of this dynamic layer can be improved by crosslinking the enzyme molecules with glutaraldehyde (Wang et al., 2008). It is important to note that there is no covalent binding Gelification on the membrane
Entrapment
Within the Retention in polymeric matrix membrane pores
Attachment on the membrane Physical adsorption
Risk of leaching Diffusional limitation
Covalent binding
Expensive and irreversible binding
Fig. 11.4 Different types of active membrane preparation and their main drawbacks.
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Membrane bioreactors and the production of food ingredients 321 between the support material and the enzymes; the membrane regeneration is thus possible by removing the enzyme gel through backwash with a highpressure water or gas. As physical entrapment in the porous support, this immobilization method is really simple and generally leads to high enzyme loading. However, in both cases all the biocatalyst molecules are not active owing to diffusion limitation or steric obstruction that hinders access of the substrate to the catalytic site of enzyme. It is not obvious to conclude which one of these two methods is the most attractive. Wang et al. (2008), who compared both immobilization methods with the same membrane reactor, observed that the performance is enhanced when the enzymes are on the surface whereas Hilal et al. (2006) reported opposite findings. Finally, enzymes can be attached by noncovalent binding (adsorption through hydrophobic or ionic interactions) or covalent binding on membrane surface. Enzyme adsorption is certainly the simplest and cheapest immobilization method because it can be achieved in one step (by immersion in enzyme solution or filtration) without the use of any activator. Even if protein adsorption is higher on hydrophobic membrane as it is generally reported and demonstrated (Shamel et al., 2007), this method can be applied to various membrane materials from hydrophilic to hydrophobic ones. Recently, Engel et al. (2008) used a strongly basic anion-exchanger membrane for immobilization of b-galactosidase in order to synthesize galacto-oligosaccharides. However, owing to the weakness of the binding force the risk of desorption is high, thus reducing the potential of such active membranes. Nevertheless, according to Tischer and Kasche (1999), immobilization via adsorption method is particularly appropriate when the membrane is used in nonaqueous solvents in which desorption phenomena may be overcome owing to the low solubility of enzymes in such solvents. Compared with noncovalent binding, covalent attachment provides virtually irreversible binding between the amino or carboxyl groups of the enzyme and functional groups of the membrane. It thus avoids biocatalyst leaching and increases enzyme stability especially in nonaqueous reaction media. However, the irreversibility of the linkage may be a serious drawback when the biocatalyst becomes inactive; both biocatalysts and support are unusable. In addition, compared with free enzymes a decrease in enzymatic activity is observed if the catalytic site is involved in the bonding reaction with the support or if the enzyme molecules are immobilized in an inactive conformation. Covalent attachment is generally considered as an expensive method of immobilization. It needs the use of chemicals (i.e. reactive groups such as carbodi-imide, cyanogen bromide, diazonium salts or reagents such like epichlorhydrin or glutaraldehyde) to form the covalent bonds and previous treatments for membrane surface activation (i.e. irradiation or chemical treatments) especially in the case of inorganic supports (Sousa et al., 2001). To overcome some of these drawbacks a new method for enzyme immobilization onto ceramic membranes has been developed at the European
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322 Separation, extraction and concentration processes Membrane Institute (IEM) (Belleville et al., 2001, Magnan et al., 2004). In order to generate functional groups on the membrane surface, the ceramic support is first coated by filtration with polymers (such as gelatine or polyethyleneimine). Then the enzymes are covalently linked to the polymeric layer by glutaraldehyde bonds. Such active membranes are of interest for several reasons linked to their inorganic nature. They can be used under pressure and with a wide range of solvents in particular in supercritical media (Gumi et al., 2007; Lozano et al., 2004). In addition, the ceramic support can be easily regenerated when the enzymes are deactivated. Furthermore, the hydrophilic nature of the polymer coating offers a good environment to preserve enzymes from deactivation by anhydrous conditions. In summary, there are many routes to prepare active membranes and as all of them have advantages and drawbacks, it is not possible to determine the best method. The choice has to be done on a case-by-case basis depending on the type of biocatalyst, membrane material and reactor configuration.
11.3 Applications of membrane bioreactors in food industries During the past two decades, the MBR market worldwide experienced a huge increase from around US$ 1 million in 1990 to US$ 296 million in 2008. It is expected to grow at an annual growth rate of 10.5% to rise to nearly US$ 0.5 billion by 2013 (Anon., 2008). However, these figures concern mainly the market of wastewater treatment driven mainly by legislation and water stress. The industrial applications of MBRs in the food area are much less widespread. Giorno (2008) who analyzed the development of MBRs in terms of publication of patents from 2004 to 2008 reported that patents in food, biotechnology and pharmaceuticals represent only 12% of the total patents whereas water treatments represent 77%. For food applications, the uses of MBRs are mainly involved in the processing of food and beverages (e.g. wine, fruit juices and milk) on the one hand and on the other hand for the production of a wide range of food ingredients obtained by biocatalysis processes (e.g. sugars, organic acids, peptides, esters). 11.3.1 Processing of food and beverages The idea of producing ethanol in a MBR is not really new. In the early 1980s, Cheryan and Mehaia reported the production of ethanol by Saccharomyces cerevisae in a MBR (Cheryan and Mehaia, 1984). They obtained MBR conversions which were 30 times higher than the one observed in batch reactor. In 1991, they associated two MBRs in series in order to produce vinegar from date juice (Mehaia and Cheryan, 1991). The sugars were first converted to ethanol by S. cerevisae in the first MBR. The permeate obtained
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Membrane bioreactors and the production of food ingredients 323 was then used to feed the second MBR in which ethanol was converted into acetic acid by Acetobacter aceti. This two-stage fermentation process exhibited a yield of about 0.5 g acetic acid per gram sugars which is about 75% of the theoretical value and close to that obtained in batch fermentations in current commercial vinegar production. In addition to its high productivity rate, the process allowed the production of a cell-free, clear product stream requiring little or no further downstream processing. More recently, Takaya et al. (2002) examined yeast growth and ethanol fermentation in MBR using a grape juice. They suggested the use of a double vessel membrane reactor, a continuous stirred tank reactor (CSTR) associated with a CRMBR, with which they observed a productivity of dry wine 28 times higher than that in the batch fermentation. Zhang and Lovitt (2006) suggested the use of MBR as a strategy for enhancing malolactic fermentation in wine and cider maturation. Concerning fruit or vegetable juices production, ultrafiltration and microfiltration represent attractive alternatives to conventional processes for juice clarification since they can be carried out in continuous mode without the need of fining agents (Álvarez et al., 2000). However, in order to improve permeate flux and reduce membrane fouling, a pretreatment of juices and pulps with pectinases is usually recommended. In general; this treatment is realized in a tank before juice filtration; however, when this reaction is carried out in a FEMBR, the product inhibition is avoided and the reactor productivity is enhanced compared with a batch system (Bélafi-Bakó et al., 2007). Rodriguez-Nogales et al. (2008), who studied the stability of such reactors for a long term experiment (15 days), observed a significant viscosity reduction (88% below the initial value) and concluded that both operations (pectin hydrolysis and juice clarification) can be achieved in a one step operation with FEMBR. In the dairy industry, the main application of MBRs and more particularly of EMBRs is lactose hydrolysis in order to improve milk digestibility. Grano et al. (2004) investigated the potentiality of a nonisothermal IEMBR. In this reactor, the feed (skimmed milk) flows along both sides of the active membrane prepared with b-galactosidase. Because of the temperature difference applied across the membrane, a thermodialysis flux occurs, from the warm to the cold side of the bioreactor and the lactose transformation occurs during the crossing. However, even if the reactor performance in terms of percentage reduction of the production time is promising and similar to those of batch and fluidized-bed reactors previously reported in the literature (Roy and Gupta, 2003), extrapolating from the laboratory-scale reactor to an industrial-scale process is not straightforward. Another type of reactor is proposed by Novalin et al. (2005). In this instance, the biocatalyst (i.e. b-galactosidase MaxilactTM, Gist-brocades) is retained in the shell side of a hollow-fibre module. Pasteurized skim milk is pumped through the hollow-fibre module (lumen side) and enzymatic solution is circulated inside the shell side. Owing to the diffusion gradient, lactose crosses the membrane
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324 Separation, extraction and concentration processes to the shell side where the reaction takes place. This diffusion reactor shows promising performances (a conversion rate around 80% within laboratoryscale operating conditions) and needs to be investigated at industrial scale. 11.3.2 Production of food ingredients The main applications of enzymatic and whole cell MBRs for food ingredients production are reported in Table 11.1 and 11.2, respectively. First, enzymatic MBRs are much more widely investigated for food ingredients production over the past 10 years than whole cell MBRs. This may be related to the difficulty of using biological cells at an industrial level, particularly owing to a decrease in activity. That may result from a nutriments limitation, product toxicity or mechanical stress, and membrane fouling which induces lower productivity. In addition, during microbial growth, many metabolites or byproducts are released in the fermentation broth at the same time as the product of interest. It is then necessary to consider further stages of separation. Nevertheless, for reactions that implicate several successive enzymes or enzyme cofactors, whole cell MBRs are more attractive than enzymatic ones. As shown in Table 11.2, whole cell MBRs are mostly investigated for the production of organic acids, or small sugars used as sweeteners. Among the various applications reported, lactic acid production is the most studied over the past few years; this is probably because of its possible use as a monomer material for the elaboration of environmentally friendly packaging material, rather than its use as a food additive. However, biosynthesis of lactic acid is an attractive process because it allows the production of high value-added products from lactose, a byproduct of cheese manufacturing. As lactic acid fermentation is characterized by product inhibition, which affects cell growth and metabolism and thus limits the production, the use of MBR is an interesting alternative. Various configurations of MBRs were investigated. Whole cell CRMBRs are the most studied despite the risk of cell denaturation owing to shear stresses. To avoid cell damage, Giorno et al. (2002) suggested working at low axial velocity (laminar flow condition); but in that instance the reactor performance is affected by fouling, which increases the hydraulic resistance of the membrane. To improve filtration performance it was suggested to use iMBR (Kamoshita et al., 1998; Schiraldi et al., 2003). Kwon et al. (2001), who considered that the main drawback of MBRs was low product concentration, suggested connecting two whole cell CRMBRs in series. They reported that 92 g L–1 lactic acid could be produced with a volumetric productivity of 57 g L–1 h–1, and suggested that this attractive value results from the continuous feeding of the second MBR with the permeate and fresh cells from the first reactor. These fresh cells are more resistant to product inhibition. To limit product inhibition it is also possible to connect the MBR to another separation process, for instance to an electrodialysis unit as suggested by Meynial-Salle et al. (2008). Nevertheless, an important disadvantage of MBRs relates to
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Table 11.1 Applications of enzymatic MBRs in agro-food industries Reaction
Objectives
Starch, Starch hydrolysis sweeteners, oligosaccharides
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Production of glucose or maltose syrups, production of maltodextrins Starch hydrolysis and Production of cyclodextrine synthesis cyclodextrins Pectin hydrolysis
Sucrose hydrolysis
Synthesis of poly-b(2-1)-fructan from sucrose Synthesis of galactooligosaccharides (GOS) from lactose
Reactor
References
CRMBR with free enzymes (amylase and/or amyloglucosidase)
Gaouar et al. (1997), Grzeskowiak-Prywecka and Słomińska (2007), Kedziora et al. (2006), Paolucci-jeanjean et al. (2000), Sarbatly and England (2004) Sakinah et al. (2008), Słomińska et al. (2002)
CRMBR with free enzymes [cyclodextrin glucosyl transferase (CGTase)] Production of pectic CRMBR with free oligosaccharides or enzymes (pectinase and/or galacturonic acids polygalacturonase) IEMBR (pectolytic enzyme preparation) Production of high CRMBR with immobilized fructose syrup enzymes (invertase adsorbed on beads) Production of IEMBR fructooligosaccharides (b-fructofuranosidase or fructosyl transferase) Production of prebiotics (GOS)
Belafi Bakó et al. (2007), Kiss et al. (2009), Olano-Martin et al. (2001) Lozano et al. (1990) Tomotani and Vitolo (2007) Hicke et al. (1999), (2006), Nishizawa et al. (2000)
CRMBR with free Czermak et al. (2004), Foda and Lopez-Leiva enzymes (b-galactosidase) (2000), Matella et al. (2006) IEMBR (b-galactosidase) Engel et al. (2008)
Membrane bioreactors and the production of food ingredients 325
Food areas
Food areas
Reaction
Objectives
Reactor
References
Proteins and peptides
Hydrolysis of whey protein
Solubilization of proteins, production of peptides with low allergenicity/or functional properties
CRMBR with free proteases (protease preparation, pepsin, chymotrypsin, carboxypeptidase, alcalase)
Cabrera-Padilla et al. (2009), Cheison et al. (2006a, 2006b), Cheison et al. (2007), Guadix et al. (2006), Mišún et al. (2008), Perea and Ugalde 1996)
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Hydrolysis of other type of proteins (casein, blood, alfafa proteins) Oil and esters
Kapel et al. (2006), Prevot-D’Alvise et al. (2004), Trusek-Holownia (2008), Wei and Chiang (2009)
Alcoholysis, acidolysis Oil engineering, esters CRMBR with free or or transesterification production immobilized lipases Monophasic IEMBR (lipase) Oil hydrolysis
Production of fatty acids, monoglycerides, diglycerides
Xu et al. (2000), Trusek-Holownia and Noworyta (2007) Giorno et al. (1997), Gumi et al. (2007), Hernández et al. (2006), Lozano et al. (2002), (2004), Pomier et al. (2005), Trusek-Holownia and Noworyta (2007) CRMBR with free lipases Gan et al. (1998) Monophasic IEMBR (lipase)
Garcia et al. (1992)
Biphasic IEMBR (lipase)
Giorno et al. (1997), Goto et al. (1992), Knezevic et al. (2004), Merçon et al. (2000), Pugazhenthi and Kumar (2004), Sachan et al. (2006), Shamel et al. (2007), Tan et al. (2002), Wang et al. (2008)
326 Separation, extraction and concentration processes
Table 11.1 Continued
Membrane bioreactors and the production of food ingredients 327 Table 11.2 Applications of whole cell MBRs in agro-food industries Objectives
Reactor
References
Production of lactic acid
CRMBR with free cells (Lactobacillus rhamnosus; Lactobacillus bulgaricus) Filtration unit spiral-sheet polymeric membrane with adsorbed cells (Bifidobacterium longum and B. helveticus) associated to a CSTR iMBR with free cells (Lactococcus lactis; Lactobacillus delbruekii)
Kwon et al. (2001), Giorno et al. (2002) Shahbazi et al. (2005)
Production of xylitol
CRMBR with free cells (Candida spp.)
Ko et al. (2008), Kwon et al. (2006)
Production of palatinose
Hollow-fibre membrane reactor with free cells immobilized in shell side (Serratia plymuthica)
Krastanov et al. (2007)
Kamoshita et al. (1998), Schiraldi et al. (2003)
Production of CRMBR with free cells (Azotobacter Cheze-Lange et al. (2002) microbial alginate vinelandii) Production of manitol
CRMBR with free cells (Leuconostoc von Weymarn et al. mesenteroides) (2002)
Production of succinic acid
CRMBR with free cells (Anaerobiospirillum succiniciproducens) associated with an electrodialysis unit
Meynial et al. (2008)
the production of metabolite or by-products with low interest. Hence, the recovery of the desired product needs further separating steps. With regard to food applications of enzymatic membrane reactors, three main areas can be distinguished: (1) starch-derived-products, sweeteners and oligosaccarides production, (2) protein hydrolysates production, (3) fatty-acid esters production and oil modification. Starch conversion into smaller assimilable sugars is the most studied reaction (Gaouar et al., 1997; Grześkowiak-Pryweck and Słomińska, 2007; Kedziora et al., 2006; Paolucci-Jeanjean et al., 2000; Sarbatly and England, 2004). Usually this reaction is carried out in a FEMBR using simultaneously amylolytic enzymes and debranching enzymes (Gaouar et al., 1997) or using liquefied starch (i.e. maltodextrin) as substrate (Grześkowiak-Pryweck and Słomińska, 2007). This was intended to produce an increased yield while preventing accumulation of limit b-dextrins in the reactor, these molecules being suspected of fouling the membrane. Paolucci-Jeanjean et al. (2000) studied the degradation of raw cassava starch into maltodextrin using Termamyl™, a thermostable a-amylase supplied by NOVO, and showed that the major drawbacks observed were the membrane fouling owing to accumulation of high molecular weight products on a one hand, and enzyme inactivation on © Woodhead Publishing Limited, 2010
328 Separation, extraction and concentration processes the other hand. In order to achieve a high level of conversion and to improve the control of system performance, Paolucci-Jeanjean et al. suggested filling the reactor with a pre-treated solution before starting the continuous feeding of the closed-loop membrane reactor with the raw starch solution. However, in a more recent critical review, Sarbatly and England (2004) presented different potential opportunities to prevent membrane fouling such as the use of particles as turbulence promoters or the use of an active membrane. In the latter case, the high molecular weight products which accumulate at the membrane surface are hydrolyzed and the degraded products can flow through the pore. Starch can also be used as a substrate for cyclodextrin production. Słomińska et al. (2002) reported that compared with the batch process, the use of MBR with cyclodextrin glucosyl transferase (CGTase) increases the process efficiency in particular when the starch concentration is high (i.e., 20% w/v). Sakinah et al. (2008) who studied the same reaction, showed that the use of gelatinized starch is unfavourable even if the enzyme tested showed a higher affinity for this substrate. Actually, the use of gelatinized starch leads to higher membrane fouling owing to the cake deposition associated with large swollen tapioca starch molecules. As reported in Table 11.1, the production of oligosaccharides can be achieved in EMBR either by polysaccharide hydrolysis (i.e. hydrolysis of pectin), or by a synthesis reaction starting from sucrose (synthesis of fructo-oligosaccharides) or lactose (synthesis of galacto-oligosaccharides). However, for fructo-oligosaccharides, it is more convenient to produce these molecules from hydrolysis of fructan-containing crop materials or directly from agave juice which is known as a natural source of fructo-oligosaccharides (Ortiz-Basurto et al., 2008). The use of enzymatic FEMBRs is also attractive to produce protein hydrolysates in particular from whey hydrolysis (Cabrera-Padilla et al., 2009; Cheison et al., 2006 a,b; Cheison et al., 2007; Guadix et al., 2006; Mišún et al., 2008; Perea and Ugalde, 1996) or from other protein sources (Kapel et al., 2006; Prevot-D’Alvise et al., 2004; Trusek-Holownia, 2008; Wei and Chiang, 2009); this is an attractive way to produce various functional and bioactive peptides. Enzymatic hydrolysis of whey proteins has been demonstrated to be an excellent method to reduce their allergenicity and the peptides obtained are widely used as food ingredients in energy-providing drinks, hypoallergenic formulae and enteral diets for children and sick adults. Although hydrolysis of protein is generally carried out in a batch reactor, this reaction can be achieved with a CRMBR. Cheison et al. (2006a, 2006b) who studied the hydrolysis of whey with protease N ‘Amano’ (Amano Enzymes Co., Nagoya, Japan) in FEMBR underlined the positive role of retentate temperature which influences the solubility of protein and the whey viscosity thus leading to higher permeate fluxes. According to them, the reactor performances depend on the flux values and enzyme concentration. At low permeate flux as well as low enzyme concentration, enzyme activity is limited
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Membrane bioreactors and the production of food ingredients 329 owing to substrate inhibition whereas high permeate flux leads to enzyme leakage. Thus, hydrodynamic properties are crucial for reactor robustness. Cabrera-Padilla et al. (2009) investigated a novel enzymatic MBR for the production of whey hydrolysates with a low content of phenylalanine. Instead of free enzymes, they chose to use carboxypeptidase A (CPA), immobilized on agarose gel particles, thus avoiding enzyme autolysis, an important drawback of processes involving proteases, and preserving the biocatalyst activity for several runs. In addition, the reactor was continuously fed with whey previously hydrolyzed with immobilized chymotrypsin to a 15.3% degree of hydrolysis. Under these conditions, they observed that their EMBR had a better performance than a conventional batch reactor associated with a diafiltration unit. Finally, the use of enzymatic MBR for the biotransformation of oils and fats has been widely reported. These bioreactors generally involve lipases or esterases which can catalyze a wide range of reactions such as hydrolysis, alcoholysis, transesterifications, aminolysis and enantiomer resolution (Hasan et al. 2006; Jaeger and Eggert, 2002). In the food industry, the major application of EMBRs is esters synthesis for the production of emulsifiers and aroma compounds and, to a lesser extent, oil hydrolysis for production of free fatty acids (FFA) and mono or diglycerides as reported in Table 11.1. Compared with the conventional process for fats and oil hydrolysis, the NaOH-catalyzed hydrolysis process, which requires high pressure of about 4.82 MPa (or more) and high temperature of about 250 °C, and even though the chemical process gives high conversion rate (97–98%), the biocatalytic route is really more attractive since it can be carried out at ambient temperature and pressure with a significant decrease of waste (Wang et al., 2008). In a first work, Garcia et al. (1992) studied the hydrolysis of milkfat in an IEMBR; the enzymes were adsorbed onto polypropylene microporous membrane and the milkfat was partially hydrolyzed during the membrane crossing. However, such a reactor leads to a mixture of free fatty acids and partially hydrolyzed milkfat. In 1998, Gan et al. (1998) investigated another type of reactor. They associated a stirred tank reactor with a de-emulsifier; the reactor previously filled with substrate (sunflower oil) and lipase was continuously fed with an aqueous buffer. The resulting emulsion was pumped through the de-emulsifier and whereas the oil phase was directly recycled in the reactor, the water phase was sent to an ultrafiltration unit. The biocatalyst, which was concentrated in the retentate, was recycled whereas glycerol was removed with the permeate. Gan et al. (1998) expected to increase the recovery of FFAs by a continuous removal of glycerol; but they observed only marginal improvements in overall reaction yield. They supposed that this was because of the unsuccessful separation of the free fatty acids produced from both the de-emulsified aqueous and oil phases. To overcome this drawback and obtain FFAs separately, most studies (Giorno et al., 1997; Goto et al., 1992; Knezevic et al., 2004; Merçon et al., 2000; Pugazhenthi and Kumar, 2004; Sachan et al., 2006; Shamel et al., 2007; Tan et al., 2002; Wang et al., 2008) refer to the use of biphasic IEMBR,
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330 Separation, extraction and concentration processes with the enzymatic membrane placed between two immiscible phases, the oil and an aqueous buffer, where the fatty acids were released. However, owing to high mass transfer resistance, the productivity of such biphasic IEMBRs is generally very low. They are thus limited to systems working in a reaction limited regime. For ester synthesis, because these reactions generally involve low molecular weight substrates having low water solubility, the use of IEMBRs with nonaqueous solvent seems to be an attractive alternative for ester bioproduction. Lozano et al. (2002) and Magnan et al. (2004) have successfully achieved the synthesis in organic solvent of butyl, butyrate and butyl laurate, respectively, using a monophasic IEMBR with a hybrid organic–inorganic catalytic membrane developed at IEM and briefly presented above. These results confirm that the hydrophilic behaviour of the inert protein ensures a proper environment for the enzymes, which are preserved from inactivation in nonaqueous media. The immobilized enzyme, (Candida antarctica lipase B (CALB)), exhibited a particularly high stability, a half-life of 202 days, (Lozano et al., 2002). Owing to the drawbacks of organic solvents and above all their toxicity, some studies used more environmentally friendly solvents such as supercritical carbon dioxide (SC CO2). Despites the numerous advantages of SC CO2, among which are the tunability of solvent properties and simple downstream processing features, only a few examples of EMBRs using SC CO2 as a reaction medium have been reported. Lozano et al. (2004) compared the butyl butyrate synthesis catalyzed by CALB in IEMBR using either SC CO2 or organic solvents (acetonitrile, acetone and hexane) as reaction media. Both activity and selectivity were enhanced in supercritical medium. However, it must be noted that the IEMBR described here was used in a batch mode without membrane permeation; the interest of the system is thus limited. Nevertheless, when the reaction is carried out in continuous dead-end filtration mode, the production of ester is enhanced compared with experiments carried out in batch and semi-batch modes as shown by Gumi et al. (2007) for butyl laurate synthesis in SC CO2. This is because the convective transfer of substrate through the membrane is much more efficient than the diffusive one. Finally, the use of EMBR for structured lipids or engineered oil production was studied by Xu et al. (2000) for the reaction between medium-chain triacylglycerols (MCT) and n-3 polyunsaturated fatty acids from fish oil using Lipozyme™ (Novo Nordisk) as a biocatalyst in a membrane reactor. They reported that the percentage incorporation of polyunsaturated fatty acids into MCT was increased by about 15% over 80 h by simultaneous separation of the released medium-chain fatty acids compared with a batch experiment. More recently, Pomier et al. (2005) investigated the enzymatic modification of castor oil through the interesterification of castor oil triglycerides and methyl oleate. The reaction catalyzed by CALB was carried out in an IEMBR and SC CO2 was used as thinning agent in order to strongly decrease oil viscosity and thus to enable membrane filtration. The stability of the immobilized
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Membrane bioreactors and the production of food ingredients 331 enzymes was checked during a 25 h period and a conversion around 30% was observed. It is worth noting that the concentration of methyl ricinoleate was higher in the permeate than in the retentate showing that reaction happened mainly in the pores of the membrane where the contact between enzymes and substrates was much more favourable than at the internal surface.
11.4 Future trends Throughout this review, the potential of MBR for food processing or food ingredients production is clearly shown. However, further investigations are needed in order to guarantee the success of these reactors at the industrial level, particularly regarding their performance, integration and potential for process intensification. Further study is required on: (1) membrane materials, (2) biocatalyst engineering, (3) process engineering. Firstly, as suggested by Meng et al. (2009) in a recent review, a comprehensive investigation should be performed to understand, control and reduce membrane fouling, and particularly to avoid severe fouling. This might lead to the development of new membrane materials or the modification of those already existing in order to obtain high transfer rates. This is particularly important to improve MBR productivity particularly for whole cell MBR. The second challenge is related to research in genetic and enzyme engineering. Investigations are required to produce at lower price enzymes showing higher activity, specificity and stability in order to increase economical interest of EMBR. Last but not least, developing new ideas and concepts as well as specific modelling tools to enable easier process integration and scale changes appears today as an holistic way to optimize the performance of MBRs and to enlarge their industrial applications. Indeed, as stressed by Drews and Kraume (2005), MBRs should be considered as hybrid reactors and not as the juxtaposition of independent bioreactor and membrane filtration units for the optimization, the modelling and the scaling-up of such processes.
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332 Separation, extraction and concentration processes Anon. (2008), ‘Membrane market is experiencing a strong growth’, Membr Technol, October 2008, 10–11. Bartling K, Thompson J U S, Pfromm P H, Czermak P and Rezac M E (2001), ‘Lipasecatalysed synthesis of geranyl acetate in n-hexane with membrane-mediated water removal’, Biotechnol Bioeng, 75, 676–681. Bélafi-Bakó K, Eszterle M, Kiss K, Nemestóthy N and Gubicza L (2007), ‘Hydrolysis of pectin by Aspergillus niger polygalacturonase in a membrane bioreactor’, J Food Eng, 78, 438–442. Belfort G (1989), ‘Membranes and bioreactors: a technical challenge in biotechnology’, Biotechnol Bioeng, 33, 1047–66. Belleville M-P, Lozano P, Iborra J L and Rios G M (2001), ‘Preparation of hybrid membranes for enzymatic reaction’, Sep Purif Technol, 25, 229–233. Bódalo A, Gómez J L, Gómez E, Máximo M F and Montiel M C (2004), ‘Study of l-aminoacylase deactivation in an ultrafiltration membrane reactor’, Enzyme Microb Technol, 35, 261–266. Cabrera-Padilla R Y, Pinto G A, Giordano R L C and Giordano R C (2009), ‘A new conception of enzymatic membrane reactor for the production of whey hydrolysates with low contents of phenylalanine’, Process Biochem, 44, 269–276. Cheison S C, Wang Z and Xu S-Y (2006a), ‘Hydrolysis of whey protein isolate in a tangential flow filter membrane reactor. I. Characterisation of permeate flux and product recovery by multivariate data analysis’, J Membr Sci, 283, 45–56. Cheison S C, Wang Z and Xu S-Y (2006b), ‘Hydrolysis of whey protein isolate in a tangential flow filter membrane reactor. II. Characterisation for the fate of the enzyme by multivariate data analysis’, J Membr Sci, 286, 322–332. Cheison S C, Wang Z and Xu S-Y (2007), ‘Use of response surface methodology to optimise the hydrolysis of whey protein isolate in a tangential flow filter membrane reactor’, J Food Eng, 80, 1134–1145. Cheryan M and Mehaia M A (1984), ‘Ethanol production in a membrane recycle bioreactor – conversion of glucose using Saccharomyces cerevisiae’, Process Biochem, 19, 204–208. Cheryan M and Mehaia M A (1986), ‘Membrane bioreactors’, In: W C McGregor, Membrane separation in biotechnology, New York, Marcel Dekker, 255–301. Chèze-Lange H, Beunard D, Dhulster P, Guillochon D, Cazé A M, Morcellet M, Saude N and Junter G-A (2002), ‘Production of microbial alginate in a membrane bioreactor’, Enzyme Microb Technol, 30, 656–661. Czermak P, Ebrahimi M, Grau K, Netz S, Sawatzki G and Pfromm P H (2004), ‘Membrane-assisted enzymatic production of galactosyl-oligosaccharides from lactose in a continuous process’, J Membr Sci, 232, 85–91. Del Amor Villa E M and Wichmann R (2005), ‘Membranes in enzymatic synthesis of biotensides from renewable sources’, Catal Today, 104, 318–322. Drews A and Kraume M (2005), ‘Process improvement by application of membrane bioreactors’, Chem Eng Res Des, 83(A3), 276–284. Engel L, Ebrahimi M and Czermak P (2008), ‘Membrane chromatography reactor system for the continuous synthesis of galactosyl-oligosaccharides’, Desalination, 224, 46–51. Foda M I and Lopez-Leiva M (2000), ‘Continuous production of oligosaccharides from whey using a membrane reactor’, Process Biochem, 35, 581–587. Gan Q, Rahmat H and Weatherley L R (1998), ‘Simultaneous reaction and separation in enzymatic hydrolysis of high oleate sunflower oil – evaluation of ultrafiltration performance and process synergy’, Chem Eng J, 71, 87–96. Gaouar O, Aymard C, Zakhia N and Rios G M (1997), ‘Enzymatic hydrolysis of cassava starch into maltose syrup in a continuous membrane reactor’, J Chem Technol Biotechnol, 69, 367–375. Garcia H S, Malcata F X, Hill C G and Amundson C H (1992), ‘Use of Candida rugosa
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334 Separation, extraction and concentration processes M, Guillochon D and Dhulster P (2006), ‘Production, in continuous enzymatic membrane reactor, of an anti-hypertensive hydrolysate from an industrial alfalfa white protein concentrate exhibiting ACE inhibitory and opioid activities’, Food Chem, 98, 120–126. Kedziora P, Le Thanh J, Lewandowicz G and Prochaska K (2006) ‘An attempt to application of continuous recycle membrane reactor for hydrolysis of oxidised derivatives of potato starch’, J Membr Sci, 282, 14–20. Kiss K, Nemestóthy N, Gubicza L and Bélafi-Bakó K (2009), ‘Vacuum assisted membrane bioreactor for enzymatic hydrolysis of pectin from various agro-wastes’, Desalination, 241, 29–33. Knezevic Z, Kukic G, Vukovic M, Bugarski B and Obradovic B (2004), ‘Operating regime of a biphasic oil/aqueous hollow-fibre reactor with immobilized lipase for oil hydrolysis’, Process Biochem, 39, 1377–1385. Ko C-H, Chiu P-C, Yang C-L and Chang K-H (2008), ‘Xylitol conversion by fermentation using five yeast strains and polyelectrolyte-assisted ultrafiltration’, Biotechnol Lett, 30, 81–86. Krastanov A, Blazheva D and Stanchev V (2007), ‘Sucrose conversion into palatinose with immobilized Serratia plymuthica cells in a hollow-fibre bioreactor’, Process Biochem, 42, 1655–1659. Kwon S, Yoo I-K, Lee W G, Chang H N and Chang Y K (2001), ‘High-rate continuous production of lactic acid by Lactobacillus rhamnosus in a two-stage membrane cellrecycle bioreactor’, Biotechnol Bioeng, 73, 25–34. Kwon S-G, Park S-W and Oh D-K (2006), ‘Increase of xylitol productivity by cellrecycle fermentation of Candida tropicalis using submerged membrane bioreactor’, J Biosci Bioeng, 101, 13–18. Long W S, Kow P C, Kamaruddin A H and Bhatia S (2005), ‘Comparison of kinetic resolution between two racemic ibuprofen esters in an enzymatic membrane reactor’, Process Biochem, 40, 2417–2425. Lozano P, Manjón A, Iborra J L, Cánovas M and Romojaro F (1990), ‘Kinetic and operational study of a cross-flow reactor with immobilized pectolytic enzymes’, Enzyme Microb Technol, 12, 499–505. Lozano P, Pérez-Marín A B, De Diego T, Gómez D, Paolucci-Jeanjean D, Belleville M-P, Rios G M and Iborra J L (2002), ‘Active membranes coated with Candida antarctica lipase B: preparation and application for continuous butyl butyrate synthesis in organic media’, J Membr Sci, 201, 55–64. Lozano P, Víllora G, Gómez D, Gayo A B, Sánchez-Conesa J A, Rubio M and Iborra J L (2004), ‘Membrane reactor with immobilized Candida antarctica lipase B for ester synthesis in supercritical carbon dioxide’, J. Supercrit Fluids, 29, 121–128. Magnan E, Catarino I, Paolucci-Jeanjean D, Preziosi-Belloy L and Belleville M-P (2004), ‘Immobilisation of lipase on a ceramic membrane: activity and stability’, J Membr Sci, 241, 161–166. Matella N J, Dolan, K D and Lee Y S (2006), Comparison of galactooligosaccharide production in free-enzyme ultrafiltration and in immobilized-enzyme systems’, J Food Sci, 71, C363–C368. Mateo C, Palomo J M, Fernandez-Lorente G, Guisan J M and Fernandez-Lafuente R (2007), ‘Improvement of enzyme activity, stability and selectivity via immobilization techniques’, Enzyme Microb Technol, 40, 1451–1463. Mehaia M A and Cheryan M (1991), ‘Fermentation of date extracts to ethanol and vinegar in batch and continuous membrane reactors’, Enzyme Microb Technol, 13, 257–261. Meng F, Chae S R, Drews A, Kraume M, Shin H-S and Yang F (2009), ‘Recent advances in membrane bioreactors (MBRs): membrane fouling and membrane material’, Water Res, 43, 1489–1512. Merçon F, Sant’Anna G L and Nobrega R (2000), ‘Enzyme hydrolysis of babassu oil in a membrane reactor’, JAOCS, 77(10), 1043–1048. © Woodhead Publishing Limited, 2010
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336 Separation, extraction and concentration processes (2008), ‘Cyclodextrin production in hollow fiber membrane reactor system: effect of substrate preparation’, Sep Purif Technol, 63, 163–171. Sanchez Marcano J and Tsotsis T T (2002), ‘Membrane bioreactors’ in Catalytic Membranes and Membrane Reactors, Weinheim, Wiley VCH, 133–168. Sarbatly R and England R (2004), ‘Critical review of membrane bioreactor system used for continuous production of hydrolyzed starch’ Chem Biochem Eng Q, 18(2), 155–165. Schiraldi C, Adduci V, Valli V, Maresca C, Giuliano M, Lamberti M, Cartenì M and De Rosa M (2003), ‘High cell density cultivation of probiotics and lactic acid production’, Biotechnol Bioeng, 82, 213–222. Shahbazi A, Mims M R, Li Y, Shirley V, Ibrahim S A and Morris A (2005), ‘Lactic acid production from cheese whey by immobilized bacteria’, Appl Biochem Biotechnol, 121–124, 529–540. Shamel M M, Ramachandran K B, Hasan M and Al-Zuhair S (2007), ‘Hydrolysis of palm and olive oils by immobilised lipase using hollow fibre reactor’, Biochem Eng J, 34, 228–235. Słomińska L, Szostek A and Grześkowiak A (2002), ‘Studies on enzymatic continuous production of cyclodextrins in an ultrafiltration membrane bioreactor’, Carbohydr Polym, 50, 423–428. Sousa H A, Rodrigues C, Klein E, Afonso C A M and Crespo J G (2001), ‘Immobilisation of pig liver esterase in hollow fibre membranes’, Enzyme Microb Technol, 29, 625–634. Takaya M, Matsumoto N and Yanase N (2002), ‘Characterization of membrane bioreactor for dry wine production’, J Biosci Bioeng, 93(2), 240–244. Tan T, Wang F and Zhang H (2002), ‘Preparation of PVA/chitosan lipase membrane reactor and its application in synthesis of monoglyceride’, J Mol Catal B: Enzym, 18, 325–331. Tischer W and Kasche V (1999), ‘Immobilized enzymes: crystals or carriers?’ Trends biotechnol, 17, 326–335. Tomotani E J and Vitolo M (2007), ‘Production of high-fructose syrup using immobilized invertase in a membrane reactor’, J Food Eng, 80, 662–667. Torras C, Nabarlatz D, Vallot G, Montané D, Garcia-Valls R (2008), ‘Composite polymeric membranes for process intensification: enzymatic hydrolysis of oligodextrans’, Chem Eng J, 144, 259–266. Trusek-Holownia A (2008), ‘Production of protein hydrolysates in an enzymatic membrane reactor’, Biochem Eng J, 39, 221–229. Trusek-Holownia A and Noworyta A (2007), ‘An integrated process: ester synthesis in an enzymatic membrane reactor and water sorption’ J Biotechnol, 130, 47–56. Valadez-Blanco R, Ferreira F C, Ferreira Jorge R and Livingston A G (2008), ‘A membrane bioreactor for biotransformations of hydrophobic molecules using organic solvent nanofiltration (OSN) membranes’, J Membr Sci, 317, 50–64. Vane L M (2005), ‘A review of pervaporation for product recovery from biomass fermentation processes’, J Chem Technol Biotechnol, 80, 603–629. von Weymarn N, Kiviharju K and Leisola M (2002), ‘High-level production of d-mannitol with membrane cell-recycle bioreactor’, J Ind Microbiol Biotechnol, 29, 44–49. Wang Y, Hu Y, Xu J, Luo G and Dai Y (2007), ‘Immobilization of lipase with a special microstructure in composite hydrophilic CA/hydrophobic PTFE membrane for the chiral separation of racemic ibuprofen’, J Membr Sci, 293, 133–141. Wang Y, Xu J, Luo G and Dai Y (2008), ‘Immobilization of lipase by ultrafiltration and cross-linking onto the polysulfone membrane surface’, Bioresour Technol, 99, 2299–2303. Wei J-T and Chiang B-H (2009), ‘Bioactive peptide production by hydrolysis of porcine blood proteins in a continuous enzymatic membrane reactor’, J Sci Food Agric, 89, 372–378.
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Membrane bioreactors and the production of food ingredients 337 Won K, Hong J-K, Kim K J and Moon S-J (2006), ‘Lipase-catalyzed enantioselective esterification of racemic ibuprofen coupled with pervaporation’, Process Biochem, 41, 264–269. Xu J, Wang Y, Hu Y, Luo G and Dai Y (2006), ‘Candida rugosa lipase immobilized by a specially designed microstructure in the PVA/PTFE composite membrane’, J Membr Sci, 281, 410–416. Xu X, Skands A, Jonsson G and Adler-Nissen J (2000), ‘Production of structured lipids by lipase-catalysed interesterification in an ultrafiltration membrane reactor’, Biotechnol Lett, 22, 1667–1671. Zhang D and Lovitt R W (2006), ‘Strategies for enhanced malolactic fermentation in wine and cider maturation’, J Chem Technol Biotechnol, 81, 1130–1140. Ziobrowski Z, Kiss K, Rotkegel A, Nemestóthy N, Krupiczka R and Gubicza L (2009), ‘Pervaporation aided enzymatic production of glycerol monostearate in organic solvents’, Desalination, 241, 212–217.
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Separation technologies in dairy and egg processing 341
12 Separation technologies in dairy and egg processing G. Gésan-Guiziou, INRA, France
Abstract: A description is given of the separation processes applied in the dairy and egg-processing sectors. The main focus is on the use of membrane and chromatography processes for protein separation, in accordance with the developed applications in these two food sectors, but an overview of all the separation processes implemented throughout the processing chains (from product reception to wastewater treatment) is also given. Characteristics of the food products are presented and the performance of the main separation techniques used on an industrial scale, in particular with respect to the operating modes and properties of the obtained fractions, is discussed. Finally, some of the options that food companies need to consider over the next decade, such as manufacture of isolated peptides and proteins fractions with specific functions and reduction of environmental impacts, are explored. Key words: dairy industry, egg-processing industry, proteins, membrane processes, chromatography.
12.1 Introduction During the past three decades, in parts of the world where food products are readily available, increasing industrial demand has been directed towards large-scale concentration, extraction and purification procedures of food components. Proteins and, to a lesser extent, lipids have been exploited not only for their basic nutritional contribution owing in particular to their essential amino acids, fatty acids and phospholipids composition, but also for their functional and biological properties which make them potential ingredients of health-promoting foods. Functional ingredients have largely been produced for modifying or enhancing the textural and rheological characteristics of © Woodhead Publishing Limited, 2010
342 Separation, extraction and concentration processes food stuffs: for instance they emulsify fat, bind and entrap water, increase the viscosity of liquids, and form gels of different characteristics. Specific food components are also associated with biological activities such as antimicrobial, antihypertensive, anticancer and opioid activities and electrolyte transfer. In the food sector, the dairy industry has undoubtedly developed the most advanced extraction/separation procedures for concentration and fractionation of molecules from milk and its derivatives. Over the last 40 years, membrane processes have become major tools for the separation of dairy components. The first membrane development in the separation procedures of milk components occurred in the late 1960s with the advent of membrane separation and since then a new industry has spawned for whey treatment as well as new avenues for cheesemaking. The separation techniques, and particularly membrane separations, have been implemented throughout the dairy processing chain: milk reception, cheesemaking, serum protein concentration, fractionation of protein, and effluents treatment. Simultaneously, the egg processing industry has seen changes, including high-speed egg-breaking machines, improved pasteurization technology and improved multistage spray driers (Stadelman and Cotteril, 1995). The egg production has sharply increased during the last 20 years (reaching 63 million tonnes of hen eggs in 2007, which corresponds to approximately one thousand billion eggs on a basis of 16 eggs per kg, Nau et al., 2010) and there has been a continuing growth of processed egg products. Today, approximately 30% of the worldwide consumption of eggs is in the form of processed egg products (Froning, 2008), approximately 25% of which is European consumption (Nau et al., 2010). Owing to their multifunctional properties, many of these egg products are used as ingredients in various food applications. But, innovative studies reveal the diversity of chemical properties, structure and function of egg components that has fuelled increasing demand to more fully utilize egg products and has led to the development of new industrial separation processes mainly based on chromatography for the extraction of components (proteins, lipids) from either the egg white or yolk. The aim of this chapter is to give a description of the separation processes applied in the dairy and egg-processing sectors. The main focus is on the use of membrane and chromatography processes for protein separation, in accordance with the developed applications in these two food sectors, but an overview of all the separation techniques implemented throughout the processing chains (from product reception to the wastewater treatment) is also given. The characteristics of the food products are presented and the performance of the main separation techniques used at industrial scale, in particular with respect to the operating modes and properties of the obtained fractions, is discussed. Finally, some of the options that food companies need to consider over the next decade, such as manufacture of isolated peptides and proteins fractions with specific functions and reduction of environmental impacts, are explored.
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Separation technologies in dairy and egg processing 343
12.2 The dairy industry and composition of dairy products It would be impossible to develop a separation process for any group of milk components without prior thorough knowledge of their physicochemical characteristics. 12.2.1 Milk Milk is defined as the secretion of the mammary glands of mammals, its primary natural function being nutrition of the young. In this chapter, the word ‘milk’ will refer to ‘normal’ milk of healthy cows. Milk is a complex fluid, with a constant pH around 6.5–6.7 at ambient temperature. A classification of the principal milk constituents is given in Table 12.1 and Fig. 12.1. The main constituents are lactose (48–50 g L–1), fat (34–44 g L–1), proteins Table 12.1 Approximate composition of cows’ milk and association state Components Concentration (g L–1) Water Fat Lactose Proteins Caseins Soluble proteins Ashes (minerals and salts)
Size (mm) or molecular Association state weight (order of magnitude) (g mol–1 or Da)
870–875 34–44 0.15–15 mm 48–50 342 Da 32–35 25–28 50–500 mm 5–7 14 200–150 000 Da
Solvent Separated phase Solution
8–9
Solution
Ions
Aggregates = micelles Mono-oligomers
Micro-organisms
Soluble proteins Casein micelles
Lactose, small organic molecules
Fat globules
Somatic cells
(mm) 0.0001
0.001
0.01
1
10
100
Microfiltration
Nanofiltration Reverse osmosis
0.1
Ultrafiltration
Fig. 12.1 Approximate particle sizes for which separation by means of membrane filtration can be applied. The size of some milk components is also indicated in comparison with membrane pore size range.
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344 Separation, extraction and concentration processes (32–35 g L–1 expressed as nitrogen, N ¥ 6.38) and minerals (8–9 g L–1), but many other small species, such as urea and vitamins, are present in the soluble phase of milk. Lactose is the distinctive carbohydrate of milk, composed of glucose and galactose. Fat is present in milk in the form of fat globules from 0.15 to 15 mm diameter. There are numerous small fat globules representing a small fraction of the fat but a large surface area, and a few large ones comprising a larger fat percentage (Mulder and Walstra, 1974). These globules are naturally surrounded by their native milk fat globule membrane, which is composed mainly of phospholipids, cholesterol, proteins, and enzymes. Proteins correspond to about 95% of the total nitrogen of the milk. Bovine milk contains numerous proteins which are classically divided into two major groups. First, the caseins, which are insoluble at their isoelectric point (pH = 4.6 at temperature >8 °C), are associated into large globular aggregates, called casein micelles. The composition and structure of the casein micelles have been studied for more than 40 years, and are still not totally elucidated (Horne, 2006; Qi, 2007). They are made of four distinct caseins, as1, as2, b, and k in the proportions 3:1:3:1, and 8% in mass of calcium and phosphate, often called the colloidal phosphate. The structural model accepted most widely has a roughly spherical, core-shell structure, with outer diameters ranging from 50 to 500 nm. The core is generally described as a homogeneous web of caseins in which calcium phosphate nanoclusters are distributed randomly. The shell is essentially made of k-caseins that extend into the aqueous phase as a polyelectrolyte brush producing steric and electrostatic repulsions between micelles. The casein micelles are ‘soft’ and ‘dynamic’ colloids. The second group are the serum proteins, also called soluble proteins because they do not precipitate in the ionic environment of milk when rennet is added or acidification occurs down to pH = 4.6. They are mainly composed of b-lactoglobulin (~ 3.2 g L–1), a-lactalbumin (~ 1.2 g L–1), bovine serum albumin, BSA (~ 0.4 g L–1), immunoglobulins (~ 0.7 g L–1), and minor proteins lactoferrin (LF) and lactoperoxidase (LP). They are largely present in milk in molecular form or as very small oligomers (Table 12.2). The mineral fraction (8–9 g L–1) contains both anions and cations (Table 12.3), which are in dynamic equilibrium among themselves in solution and between solution and proteins (Fig. 12.2). Changing the external conditions of milk, such as pH and temperature, may cause alterations in the mineral equilibriums that could induce modifications in the structure and stability of casein micelles (owing to colloidal phosphate), and then changes in the performance of fractionation processes. Milk collected by the dairy plant also contains a microbial flora formed by numerous species illustrating the contaminations of milk by the udder, the milking machine, the local farm atmosphere, the bulk tank and the transportation equipment. These bacteria account for about 0.01% of the volume of milk of healthy cows. Somatic cells, coming from the bovine © Woodhead Publishing Limited, 2010
Separation technologies in dairy and egg processing 345 Table 12.2 Characteristics (concentration, isoelectric point, molecular weight) and potential biological functions of some proteins found in milk Concentration Isoelectric Molecular Biological (g L–1) point weight (kg functions* mol–1 or kDa) Caseins 25–28 ≈4.6 14–22 as1, as2, k, b
Iron carrier, immunomodulator, precursor for bioactive peptides
b-lactoglobulin 3.2 5.4 18.4 (dimer in milk 36.8)
Retinol carrier, potential antioxidant, precursor for bioactive peptides, binds fatty acids
a-lactalbumin 1.2 4.4 14.2
Lactose synthesis in mammary gland, calcium carrier, immunomodulator, precursor for bioactive peptide
Immunoglobulins 0.5–1.0 5-8 150–1000
Specific immune protection (antibodies and complement system), potential precursor for bioactive peptides
Bovine serum 0.4 5.1 66.3 albumin
Precursors of bioactive peptides
Lactoferrin 0.2 7.9 80
Antimicrobial, antioxidative, anticarcinogenic, anti-inflammatory, iron transport, cell growth, regulation, precursor for bioactive peptides, immunomodulator
Lactoperoxidase 0.03 9.6 78
Antimicrobial, synergetic effect with immunoglobulins and lactoferrin
*Adapted from Korhonen and Pihlanto (2007).
mammary gland (leucocytes, macrophage and epithelial cells) are also normally present in milk (100 000–400 000 cells mL–1). One particular property of milk is that it is not stable. Several changes occur as a result of physical and chemical changes, for instance when temperature
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346 Separation, extraction and concentration processes Table 12.3 Mineral composition of cows’ milk Component
Total concentration in milk (mg kg–1)
Concentration in the soluble phase (mg kg–1)
Calcium, Ca Magnesium, Mg Sodium, Na Potassium, K Chloride, Cl Phosphorus, P Non-colloidal phosphate (expressed in P) Citrate
1250 115 425 1600 1100 950 720
350 70 400 1500 1100 420 300
1650
1500
Non-protein fraction CaCit–
CaHPO4 3b
2b
3a
2a
1a
Ca2+
1b
Colloidal fraction (casein micelles)
HPO42–
Cit3–
H+ 5b HCit2–
5a
4a
4b H2PO4–
Fig. 12.2 Physicochemical equilibria between milk components; Cit, citrate (adapted from Gaucheron, 2005).
is decreased (creaming of the fat owing to aggregation of the globules, variation in the salt composition, for instance), or owing to biochemical changes induced by active endogenous enzymes (proteases and lipases, for instance) and micro-organisms contained in milk, the best effect being the degradation of lactose and production of lactic acid causing a decrease in pH. 12.2.2 Whey Whey is the liquid co-product of cheese-making. At a first glance, it can be considered as milk without casein micelles and fat. Whey contains approximately 65 g L–1 dry matter: mainly lactose (~50 g L–1), nitrogen © Woodhead Publishing Limited, 2010
Separation technologies in dairy and egg processing 347 matter (proteins ~6 g L–1), ashes (minerals ~6 g L–1) and fat (0.3 g L–1) (Table 12.4). The composition of whey varies depending on the type of cheese produced and, more particularly, on the methods used for the coagulation of the milk. There are two main types of whey (Table 12.4): the ‘sweet’ whey is obtained by enzymatic hydrolysis of casein using rennet and its final pH is ~6.0–6.6; it contains the soluble glycomacropeptide portion of the k-casein that is released in the serum phase under the action of chymosin. The ‘acid’ whey is obtained after biological or chemical acidification of milk down to the isoelectric point of the caseins (pH 4.6) (Table 12.4). The acid whey has a higher mineral content than the sweet whey because of the release of minerals from the casein micelles (mainly calcium and phosphate) into the serum phase under acid conditions (Table 12.4, Fig. 12.2). Despite the complex composition of milk and whey, the main dairy constituents (fat, casein micelles, serum proteins, lactose and minerals) are well separated from others in terms of size (Fig. 12.1). The dairy fluids are therefore ideal starting materials for membrane separation processes although other industrial separation processes are developed on the basis of specific characteristics of milk constituents, such as differences in density, coagulation abilities and charge of constituents.
12.3 Pretreatment of milk using separation techniques 12.3.1 Skimming of milk and fractionation of fat globules Because manufacturers set a standard for their product and because most extraction procedures for milk proteins use skimmed milk as the starting fluid, Table 12.4 Approximate composition of sweet and acid wheys (from Marshall, 1982)
Sweet whey (g L–1)
Acid whey (g L–1)
Dry matter Nitrogen matter Non-nitrogen matter Lactose Ashes Fat Calcium Sulfate Magnesium Sodium Potassium Chloride lactate pH
66 6.2 0.37 52.3 5.0 0.2 0.5 0.7 0.07 0.53 1.45 1.02 – 6.4
64 5.8 0.40 44.3 7.5 0.3 1.6 0.5 0.10 0.51 1.40 0.90 6.4 4.6
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348 Separation, extraction and concentration processes standardization of the fat content of milk is classically applied in industrial manufacture. Separation of cream from skimmed milk has been a common practice for over 100 years, and is usually performed, at around 50 °C, by means of a flow-through centrifuge called a cream separator. As milk fat has a lower density than plasma, the fat globules rise under the influence of gravity, and the rate of rising is increased when a centrifugal field is applied and when the centrifuge is equipped with conical discs so as to limit the distance over which the fat globules have to move. Despite its widespread application in the dairy industry, this separation is not totally perfect and the skimmed milk, which generally represents 90% of the volume of the entering whole milk, still contains a very low fat content, 0.05 to 0.08%, which can strongly influence the efficiency of downstream processes and the resulting properties of dairy products. Moreover, this technology can not separate the native milk fat globules according to their sizes. In many dairy products, fat composition and structure cannot be adjusted easily because fat is usually homogenized. Homogenization reduces the fat globule size that results in small fat droplets that are disrupted and covered by caseins and serum proteins. Such modification of the fat globules characteristics promote their interactions with the cheese casein matrix which affect the structure of the rennet gel and decrease product functionality. Recently, separation of milk fat into small and large globules was proposed using ceramic microfiltration (MF) membranes under hydrodynamic conditions, which causes no damage to the native fat globule membrane (Goudédranche et al., 2000; Michalski et al., 2006). The process feasibility was shown but the development of industrial applications depends on the profits associated with the new products. It is claimed that the use of the small globule fraction in cheese production gives a smoother and finer texture, probably because of the interaction of fat globule membrane with the cheese casein matrix, and the differences in triglycerides content of the fat globules according to their size (Michalski et al., 2007). 12.3.2 Bacterial removal Centrifugation (CF) is also used to separate particles that have a density larger than that of milk serum, such as dirt particles, somatic cells, and micro-organisms. The bacterial quality of milk is the most variable factor with which the manufacturer has to contend. Pathogenic bacteria can contaminate milk and a wide variety of micro-organisms can lead to various cheese defects. Thus, in order to minimize the health hazard and control bacterial growth during milk processing, various combinations of time–temperature treatments (such as pasteurization 72 °C for 12 s) can be used, but they almost always affect the flavour, the functionalities of the milk components, and the cheesemaking properties. CF and crossflow MF are non-thermal preservation technologies that have been widely applied for removal of bacteria (Gésan-Guiziou, 2010).
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Separation technologies in dairy and egg processing 349 CF is mainly used for the removal of bacteria and spores from minimally pasteurized milk. It is especially applicable for spores removal, as their density is higher than that of bacteria, despite their smaller size. Its most important application has therefore proven to be for the removal of spores that undergo a late acid fermentation in the semi-hard and hard cheese process. It is also used to produce extended-shelf-life pasteurized milk, with a gain of about 3–5 days. Centrifugation is often referred as ‘bactofugation’ because the major manufacturer (Tetra Pak) produces centrifuges under the trademark of Bactofuge®. The separation is classically performed at 50–60 °C, which is a temperature similar to that in the cream separator, as the bactofuge is normally installed in series with the cream separator. At this temperature, the total bacteria count is reduced by 86.0–92.0% (decimal reduction of ~ 1 log), the spore removal reaches 90.0–98.0% (decimal reduction ranging from 1.0 up to 1.7), and a significant protein loss is observed, reaching 2.5 to 12% according to the types of machines. To reduce effectively the load of bacterial spores, a double bactofugation is practised in the cheese industry. An alternative to bactofugation is the use of MF, which was introduced successfully in the 1990s (Gésan-Guiziou, 2010). MF is particularly adapted to the removal of bacteria from skimmed milk, because the size ranges of fat globules and bacteria overlap (Fig. 12.1). The MF of whole milk would lead to a contaminated fat fraction which could be difficult to valorize. This separation is classically performed with a multichannel ceramic membrane with a pore diameter of 1.4 mm at a temperature of 35–55 °C: the milk components permeate through the membrane whereas the bacteria are satisfactorily retained. In order to overcome the membrane fouling phenomena, most MF plants in dairies operate using the hydraulic concept of uniform transmembrane pressure (UTP), that involves the circulation of the permeate co-current of the retentate (Alva-Laval Company, Bactocatch system) (Fig. 12.3) (Sandblom, 1974). In order to create a large pressure drop on the permeate side, this compartment is filled with plastic balls (Bactocatch system, Alva-Laval) or membranes are placed into small stainless-steel tubes in order to reduce the external space between the housing and the membrane porous media (Invensys APV). More recently, ceramic membranes with linear hydraulic resistance gradient were commercialized for this application: GP Membralox® membranes from Pall-Exekia (GP for permeability gradient); and Isoflux® membranes from Tami-Industries. This new membrane concept, with higher hydraulic resistance at the entrance of the membrane when the transmembrane pressure (TMP) is high, creates homogeneous filtration performance all along the membrane without a permeate circulation loop. By combining a high crossflow velocity (6–9 m s–1) and low TMP (~50 kPa), permeation fluxes are high (400–650 L h–1 m–2) for ~10 h, and matter losses are low: ~5% of the entering milk volume with a volume reduction factor (VRF), the ratio between either the volumes of feed and retentate in discontinuous mode, or the flow-rates of feed and retentate in continuous
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350 Separation, extraction and concentration processes
Permeate P
P
Retentate
Product
P
P
Fig. 12.3 Principles of the uniform transmembrane pressure system (Sandblom, 1974): the permeate circulates in co-current with the retentate in order to create a constant transmembrane pressure all along the filtering path. P, pressure sensor. Raw whole milk Skimming by centrifugation 50–60 °C
Cream
Skimmed milk 1.4 mm microfiltration 35–55 °C, VRF = 20
Retentate 1.4 mm microfiltration 35–55 °C, VRF = 10
Heat treatment
Heat-treated cream
Retentate
Microfiltered skimmed milk
Mixing, homogenization
Microfiltered whole milk
Fig. 12.4 Schematic representation of process for microfiltration of whole milk. Dashed lines, options for the treatment of retentate. VRF, volume reduction factor.
mode of 20, and ~0.5% with VRF = 200 (Fig. 12.4). The residual fat is largely retained (63%) and the proteins mostly transmitted (99%). The decimal reduction in bacteria count reaches 3–4 log with a high decimal reduction in spore count (2–4 log) that can be attributed to the binding of spores with cell walls. Somatic cells are also totally retained by the 1.4 mm
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Separation technologies in dairy and egg processing 351 MF membrane and, thus, the microfiltered milk is not degraded by their thermostable enzymes. MF is then much more efficient than bactofugation and makes it possible to decrease the microbial load of milk, while maintaining the organoleptic quality of milk owing to heat treatment at low temperatures (37–50 °C). The fat content of the microfiltered milk can be adjusted by addition of heat-treated cream (Fig. 12.4). The final retentate, that contains the bacteria and somatic cells in the milk, can be discharged separately for other suitable applications; blended with the cream and similarly heat-treated; or fed back to the cream for repeated separation. This process has been commercialized either for drinking or cheese milks. In most countries, mainly in Europe and North America (in particular Canada), the microfiltered milk intended for the drinking market undergoes a final pasteurization step to meet current regulatory requirements and lead to an extended life of 20–32 days at 4–6 °C. In cheesemaking, this process is used to produce safe raw milk cheeses, but requires knowledge about the microbial ecosystem composition that should be added to the treated milk for obtaining typical ripening of cheeses.
12.4 Standardization and concentration of milk proteins in the dairy industry The first milk protein concentrates were obtained by complexation of serum proteins with caseins after heat denaturation at ~90–95 °C (1–20 min), acid conditions (pH ~5.8–4.6) and possible addition of CaCl2. Precipitated products, referred to as ‘casein-whey protein co-precipitates’, were recovered with high yield (92–95%) reaching 97% with addition of calcium. However, after redispersion of the precipitates in the presence of a Ca chelatant (tripolyphosphate), proteins produced by this method had poor solubility properties. In the late 1960s, the development of membrane technology and more particularly of ultrafiltration revolutionized the dairy industry, offering the possibility of concentrating undenatured proteins. An ultrafiltration (UF) step was directly incorporated into the cheesemaking in what is known as the MMV process after the investigators Maubois, Mocquot and Vassal, who originally developed this process in 1969 (Maubois, 1981). Currently, UF is the most widely used membrane process for total protein concentration in cheese manufacture (Mistry and Maubois, 2004). The general idea behind this process is to concentrate all the milk proteins and simultaneously to remove the excess water and lactose by an initial UF step before coagulation, induced by addition of enzymes (rennet) and micro-organisms, thereby reducing or eliminating the need to separate the whey from the curd (Fig. 12.5). Spiral wound polymer membranes and to a lesser extent tubular ceramic membranes with a cut-off 10–50 kg mol–1 are used for this application, at a transmembrane pressure of 200–400 kPa and flux of 30–120 L h–1 m–2. © Woodhead Publishing Limited, 2010
352 Separation, extraction and concentration processes Traditional method
Ultrafiltration, MMV process
Milk
Milk Ultrafiltration
Rennet starters Coagulation
Pre-cheese Rennet starters
Coagulum
Draining
Cheese
Whey
Ultrafiltrate
Cheese
Fig. 12.5 Schematic representations of traditional ways of making cheese and the MMV process (with concentration to the final protein composition of the cheese).
Three categories of protein concentration level obtained by UF can be distinguished: standardization (retentates with VRF up to 1.7), intermediate concentration (1.7 < VRF < 5) and total concentration (VRF ~ 5–7). For almost 40 years, UF has been used in milk-processing plants, for total protein standardization (VRF up to 1.7) of both drinking and cheese milks. The main advantages of protein standardization, which results in a constant milk composition all year round independent of seasonal variations, are a more efficient use of the machinery and a better process control of the continuous cheesemaking processes. In addition, these concentrations provide economic benefits in terms of reduced requirement for coagulating enzymes and starters, as well as a slight increase in cheese yield, attributed to reduced losses of fat and casein particles in whey and better retention of serum proteins in the aqueous phase of cheese (Mistry and Maubois, 2004). An indirect but serious advantage of UF, admitted by the Codex Alimentarius, is the use of permeate to decrease the protein content of drinking milk to the minimum required value by law, 28 g L–1 (2.8% w/w) in most European countries. One can note that UF with VRF ~ 1.5 can also be incorporated into the manufacture of yoghurt and related fermented products, such as koumiss (Tamine and Robinson, 2007) in order to increase the protein levels without addition of milk powders. This technology is preferred to reverse osmosis because of the reduced level of lactose in the milk base. UF results in a better yoghurt gel strength compared with those obtained with addition of skimmed
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Separation technologies in dairy and egg processing 353 milk powders or concentration by vacuum evaporation. However, compared with cheesemaking, the increase of protein level is not accompanied in the yoghurt manufacture by a yield improvement, because the high-temperature heat treatment generally applied to the milk (typically 80 °C for 10 min) leads to the attachment of denaturated serum proteins with casein micelles, resulting in the concentration of serum proteins with casein micelles after heat treatment. Processing UF at higher concentration (VRF > 1.7) leads to a further advantage in terms of cheese yield. A distinction must be made between retentates concentrated to an intermediate level (1.7 < VRF < 5) in which some serum proteins are retained and syneresis still takes place, and concentration to the final composition of the cheese (VRF ~5-7), also called pre-cheeses, where the protein content is similar to the protein content of the final cheese and very little whey drainage is observed (Fig. 12.5). Intermediate concentrated retentates have been applied to numerous cheese varieties, ranging from soft to hard. Feta cheese manufacture, which is the main industrial application of this operation, yield increases of 14% while maintaining good quality products. The pre-cheeses concept was first applied to camembert cheese but many applications have been developed successfully for the manufacture of other cheese varieties, mainly fresh unripened cheeses such as quarg, ricotta and cream cheese, and soft and semi-hard cheeses including mozarella, Saint Paulin and feta, the manufacture of which being unquestionably the greatest success worldwide of the MMV process. The advantages of the MMV process are numerous. The overall cheese yield is about 10–30% higher than in the traditional process owing mainly to the retention of serum proteins and enzyme usage is generally reduced. The MMV process eliminates the need for the large storage tanks traditionally used for heating and cooking the curds, resulting in a saving in both capital investment and energy costs. In addition, it is possible to use the MMV process to convert much of the cheese production in a continuous operation, leading to significant advantages in terms of overall capital costs and operational efficiency. However, during UF, the concentration of serum proteins, with higher water-binding capacities than caseins, and of minerals associated with the casein micelles, that are lost in the whey during coagulation in a conventional process, led to the inherent properties of UF milk cheeses. Thus, rather than duplicating traditional cheese varieties, which requires numerous adaptations (Mistry and Maubois, 2004), UF technology leads to the development of new types of cheeses, well-accepted by consumers. For fifteen years, manufacturing technologies including UF have also emerged and have become important for the production of milk protein concentrates, presenting interesting new technical possibilities in cheesemaking. These concentrates contain both major milk protein groups in proportions similar to milk and are often dried. Because there are no specific standards
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354 Separation, extraction and concentration processes of identity for such concentrates, they cover a wide range of compositional and functional characteristics and are manufactured mainly using UF and diafiltration, for the reduction of the lactose content. Milk MF has also recently emerged for the production of protein concentrates with various casein–serum proteins ratio (12.5.2).
12.5 Isolation of whole casein in the dairy industry The characteristics of the two groups of milk proteins (casein and serum proteins) differ significantly, permitting their ready separation from each other by at least three techniques. 12.5.1 Isoelectric precipitation and rennet coagulation There are two principal ways of manufacturing casein on industrial scale: isoelectric precipitation and rennet coagulation, both based on the precipitation/ aggregation of casein. In isoelectric precipitation, destabilization of the casein micelles can be accomplished at pH around 4.6: when milk is acidified, the net charge of the micelles decreases, the calcium and phosphate are removed and the micelles become less and less stable until caseins precipitate. Acidification of the milk may be carried out by one of the following processes: – Inoculation of milk with a mixed or multiple defined strain starters, such as lactic acid-producing bacteria, which degrade some of the lactose to lactic acid during the period of incubation (about 14–18 h). – Direct addition of dilute mineral acid to the milk. Adding acid (HCl or H2SO4) to milk heated at 50 °C results in a casein curd (‘acid casein’) that can be separated by straining and washing several times. The coproduct, which is a mixture of acid whey and part of the washing water, has a high ash and chloride content (where HCl is used), which poses some difficulties during spray-drying. Acid destabilization of casein is exploited in the manufacture of cottage cheese, quarg and fermented milks. – Indirect acidification of the milk. A number of alternative processes have been proposed and patented to improve the quality of the acid whey. The acidification can be performed by contact of all or part of the milk with cation-exchange resins (strong hydrogen ion exchange resin), electrodialysis (Noël, 1992) or using carbon dioxide (Hofland et al., 2003). It is believed that these processes have relatively low commercial significance (Maubois and Ollivier, 1997). Regardless of the processes used, isoelectric caseins may be dried and used as such but they are insoluble in water, which limits their applications. Therefore caseins are usually converted to a soluble form, namely caseinate, © Woodhead Publishing Limited, 2010
Separation technologies in dairy and egg processing 355 by neutralization with alkali, usually NaOH (producing sodium caseinate) and less frequently KOH, NH4OH or CaOH2. Soluble caseinates may then be sprayed or roller dried. Destabilization of casein can also be caused by the proteolysis of the micelle stabilizing protein, namely k-casein. The chymosin, the principal enzyme present in calf rennet, releases the glycomacropeptide (C-terminal segments of glycosylated k-casein), and renders the casein micelles susceptible to precipitation with Ca2+ at the natural concentration in milk. This principle is used in the manufacture of rennet casein and most cheese varieties. 12.5.2 Microfiltration for fractionation of casein micelles and serum proteins The most promising technology for the selective separation of casein micelles is undoubtedly membrane MF using a membrane with a pore size of 0.1–0.2 mm. For the past 15 years, this operation has enjoyed rapid industrial development and is still expected to increase significantly over the next few years owing to the high quality and properties of the two fractions produced: – The retentate, enriched specifically in native casein micelles, improves the rennet coagulability in the cheesemaking process. Such casein enrichment reduces rennet coagulation time, accelerates curd firmness kinetics and increases final curd firmness compared with milk or caseinate. Consequently, casein and fat retentions into the cheese curd are significantly improved (less fines and fat in the drained whey) and that leads to an increase in cheese yield. The fact that a smaller amount of serum proteins end up in the cheese manufacturing process is also advantageous, because there are then fewer flavour and texture defects attributed to serum proteins and also fewer detrimental effects of heat treatment on rennet milk coagulability. The decrease in b-lactoglobulin (b-LG) content evidently lowers the extent of the formation of the complex k-casein/b-LG during heat treatment and, consequently, renneted micelles aggregate well. This property was used to develop a new high-heat milk powder, with a cheesemaking ability similar to that of raw milk (Fig. 12.6, Quiblier et al., 1991). – The permeate, with a composition close to that of a whey, contains soluble proteins in their native state, is sterile, and free of phage particles, cellular debris, fat and glycomacropeptide. It therefore becomes a useful fluid for preparing whey protein concentrates (WPC) and isolates (WPI) (protein ratio up to 97%) with very high functional and nutritional properties, in particular compared with WPC obtained from cheese whey (Foegeding et al., 2009). Despite these advantages, the microfiltrate is sometimes used in the cheesemaking process by reincorporation in the caseinenriched fraction after denaturation under moderate concentration and heat treatment (Fig. 12.6).
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356 Separation, extraction and concentration processes Skimmed milk MF 0.1 mm VRF = 3 Microfiltrate
Retentate Cream
UF 10–20 kg mol–1 VRF = 6
Cheese
Permeate
Concentration spray drying HT (95 °C, few minutes) Retentate WPC
Milk powder b-Lactoglobulin High cheese-making ability
Fig. 12.6 Primin process (Quiblier et al., 1991). VRF, volume reduction factor; WPC, whey protein concentrates; HT, heat treatment.
Finally, purified casein micelles obtained by diafiltration of retentate against water, and WPI obtained after UF of milk microfiltrate, are excellent starting fluids for further fractionation of individual caseins or serum proteins. From an engineering point of view, this separation is classically conducted using ceramic membrane either with the UTP system or membranes with linear hydraulic resistance gradient, as is done for bacterial removal (12.3.2). Separation is usually performed at 50–55 °C, at a crossflow velocity of 7 m s–1, TMP ~ 50 kPa and VRF of 2–4. At VRF = 3, serum proteins transmission is 65–80% depending on the milk heat treatment (pasteurization, thermalization or raw skimmed milk). Permeation flux is ª75 L h–1 m–2 for 10 h according to the critical stability criterion, which allows the industry to perform long runs with very moderate fouling and high selectivity (Gésan-Guiziou et al., 1999). Owing to the high running costs and investment required by the tubular ceramic equipment, the industry is starting to operate MF with spiral-wound polymer membranes. These membranes, which can not be used with the UTP system, are operated at low temperature (<10 °C) to avoid bacterial growth. Permeation flux is reported to be very low <10 L h–1 m–2 with a low protein transmission of 10–20%. This high serum protein rejection requires diafiltration against ultrafiltrate to decrease the soluble protein level in the enriched casein fraction. According to the manufacturers, about 25% of MF plants are equipped with organic membranes today and this proportion is likely to increase in the next years for cheese applications.
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Separation technologies in dairy and egg processing 357
12.6 Separation techniques applied to whey and derivatives in the production of cheese Because of the high water content and relatively high sugar and mineral contents, more than two thirds of whey production was disposed of as waste in the 1980s with a detrimental effect on the environment. Today, a significant amount of the whey produced is being processed, mainly for the recovery of its protein content. 12.6.1 Concentration of serum proteins It has long been the practice to obtain fully denaturated proteins by heating acidified whey (temperature >>90 °C and pH << 6). This method, derived from the technology used for making whey cheeses such as ricotta in Italy, results in a precipitate containing ~80% of the initial whey proteins. After centrifugation and spray drying, this protein preparation, called ‘lactalbumin’, was still impure and because of its high insolubility in water and very poor functionality found limited use (Pearce, 1992). The availability of UF membrane processes in the 1970s offered new possibilities to fully exploit the interesting nutritional, biological, functional and, in particular, solubility properties of the serum proteins. Whey has thus become to be regarded as a starting material for processing. Currently, two main processes have been industrially designed to produce whey protein concentrates, WPC containing 35–80% protein (expressed in nitrogen N ¥ 6.38 over dry matter), and whey protein isolates, WPI containing ≥ 90%. Most whey protein concentrates are produced using membrane technologies, and the most successful application of membrane processes in the dairy industry is the production of WPC by UF. UF membranes, with an appropriate cut-off (~10–20 kg mol–1) are used to remove both the lactose and ions, yielding a retentate with a high protein concentration, which can be further processed by evaporation and spray drying. The protein content of the final whey protein product depends on the degree of concentration during UF. For 35% WPC, a VRF of 4.5–7.0 is required, although it should reach 13–20 for 50–60% WPC. Combined with a diafiltration, which removes minerals and lactose from the retentate, whey UF (VRF 30–35) can lead to WPC purity of 75–85%. During UF, membrane fouling is mainly attributed to three different species: presence of residual lipids coming from the cheese manufacture, precipitation of calcium phosphate enhanced under neutral pH (7.0–7.5) and high temperature (55 °C) and accumulation of proteins at the membrane surface, more pronounced at pH close to their isolectric point (pH~ 5.0–5.5). Several whey pre-treatments, some using membrane operations, have been proposed to increase the purity of the final concentrates (particularly by reducing the residual lipids content) and to improve UF performance (in particular by limiting calcium phosphate precipitation and protein accumulation) (Maubois and Ollivier, 1997). Similar to WPC by
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358 Separation, extraction and concentration processes ultrafiltration, WPI can be produced using the microfiltration permeate of skimmed milk. In industrial processes, whey UF is performed mainly in multistage spiralwound systems with polyethersulfone membranes. Processes are currently operated at ª50 °C, requiring a pre-treatment in order to avoid severe fouling during operation. Flux is about twice as a high as flux at 10 °C, which is a major incentive for operating at such a high temperature. However, owing to the rapid decrease in membrane prices, a low temperature process (<15 °C) is now favoured to maintain the microbiological quality of the end product. Because serum proteins are amphoteric molecules, highly purified whey protein concentrates (WPI) are produced using ion-exchange chromatography, which provides an additional level of selectivity above membrane processing. To date, ion exchange is the main adsorption technique employed, but, with the advent of new technologies, there are potential applications of affinity binding and other techniques that have historically been economically limiting. At pH lower than their isoelectric point (pH ~5.0–5.5) proteins become positive and they can be adsorbed on cation exchangers, and at pH above they can be adsorbed on anion exchangers. After the adsorption of the proteins on the resins or columns (generally performed at pH ~3.2), lactose and other non-protein whey components are rinsed with water from the ion exchanger. Then alkali is added to a pH of about 8 to desorb proteins, and the desorbed proteins are eluted from the ion exchanger, concentrated by UF, evaporated and spray-dried. There are several ways of producing protein isolates, using, in particular, specific conditions of pH, ionic strength and ionic nature, which give the processors the opportunity to adjust the composition and functionality of the obtained fractions. Among them, two major ion exchangers are commercially available for this application: the Vistec and Spherosil processes. The Vistec process uses a cellulose-based cation exchanger, in a stirred tank reactor. WPI are then produced by a single-stage batch capture of proteins, whereas the Spherosil process uses a fixed-bed ion exchanger with porous silica beads coated with a polymer material having either cationic exchange potential (–SO3H groups, Spherosil S) or anionic exchange reactivity (–N(CH3)3 groups, Spherosil QMA). Acidified whey (pH <4.6) is generally applied to Spherosil S and sweet whey (pH 5.5) to Spherosil QMA. Using chromatography, protein concentrates are always devoid of lipids, which have high foaming properties. However, UF operation is, for economical reasons, the preferred method of processing. 12.6.2 Concentration and demineralization of whey and derivatives Concentration of whey and of various ultrafiltrates at their production site is the major application of reverse osmosis (RO) owing to its flexibility and energy consumption (9–20 kWh m–3 water removed) compared with vacuum evaporation (ª100 kWh m–3). © Woodhead Publishing Limited, 2010
Separation technologies in dairy and egg processing 359 RO, which is applied to remove water, is rarely applied to milk because the flux is small, the maximum concentration attainable is low, and this technology is not as attractive as UF in cheese and yoghurt manufacturing. However, RO of milk is sometimes used as a first step in the manufacture of milk powder to increase the capacity of evaporation plants and to reduce transport costs, as is commonly done for whey RO. RO is also used to concentrate ultrafiltrates, containing about 5% total solids, mainly lactose, salt and other minor soluble components of the milk. Concentrated ultrafiltrates have valuable uses such as animal feed, recovery of lactose after crystallization, fermentation of lactose to glucose and galactose as sweetener for confectionery industry, alcohol, and lactic acid. Concentration of whey by RO to a VRF of 4 and, hence, to about 25–28% dry matter is possible. Concentration is limited by high osmotic pressure, high retentate viscosity, calcium phosphate precipitation and lactose crystallization. The RO permeate can be reused for preparing cleaning solutions, but its composition is not similar to pure water. Urea can pass to some extent, and some salts and low-molar-mass peptides may do so. Because of the high salt content of whey, which generates numerous processing difficulties and nutritional imbalance (particularly in infant food), it becomes advantageous to demineralize whey before evaporation. The demineralization of whey can be achieved in various ways (electrodialysis, ion exchange, nanofiltration) according to the type of treated whey and the required demineralization level (Table 12.5). Regardless of the chosen processes, ions, and not undissociated salts, are removed. However, for a given overall proportion of salts removed, the rate of removal varies with the kind of ions and with the technology used. Nanofiltration (NF) for instance removes monovalent ions (such as Cl–) and concentrates divalent nutrition value ions such as calcium with the proteins. Whey is currently demineralized in the range 50–95% by electrodialysis and/or ion exchange, after being concentrated by RO, but these operations lead to large volumes of polluting effluents and high investment and running costs. NF is less efficient for salt removal than electrodialysis (which
Table 12.5 Demineralization of whey. Optimal processes as a function of whey type and demineralization level required
Demineralization level (% dry matter)
30%
50–70%
Sweet whey Nanofiltration Ion exchange + (6% dry matter) nanofiltration Concentrated Electrodialysis Electrodialysis whey (18–24% dry matter)
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>90% Electrodialysis + ion exchange + nanofiltration Ion exchange + electrodialysis
360 Separation, extraction and concentration processes can achieve 70% salt removal), but has the advantage of simultaneously concentrating the liquid, which is often desired, with low losses of lactose and nitrogen. Owing to the low osmotic pressure difference between retentate and permeate compared with RO (attributed to the transfer of monovalent ions), the TMP is lower and the operation is generally more cost-effective than electrodialysis. In addition, the NF step significantly improves the technological characteristics of the concentrate and gives it higher value (increase in yield of lactose crystallization, and low hygroscopy of obtained powders). In 10 years, nanofiltration has become the industrial method of choice for partial desalting of whey. It makes it possible to reach simultaneously the concentration of dry matter (20–22% at VRF ~4) and demineralization (25–50% and even 90% with diafiltration).
12.7 Fractionation of individual proteins and peptides in the dairy industry To a large extent, the properties of the WPCs and WPIs approximate the properties of b-LG, because it constitutes more than half of the whey proteins. To exploit the particular properties of individual proteins, which are known to exert a wide range of nutritional, functional and biological activities (Tables 12.2 and 12.6), and to emphasize the properties of b-LG, fractionation of Table 12.6 Milk protein functionality in foods (Maubois and Ollivier, 1997) Functional property
Mode of action
Food system
Solubility
Protein solvation
Beverages
Water adsorption and binding
Hydrogen bonding of water, entrapment of water
Meat sausages, cakes, bread
Viscosity Thickening, water binding
Soups, gravy, salad dressing
Gelation
Protein matrix formation and setting
Meats, curds, baked goods, cheese
Cohesion–adhesion
Protein acts as adhesive material
Meat sausages, baked goods, pasta products
Elasticity
Hydrophobic bonding in gluten, disulfide links in gels
Meats, bakery
Emulsification Formation and stabilization of fat emulsions
Sausages, salad dressing, coffee whitener, soup, cakes, infant formula
Fat absorption
Binding of free fat
Sausages, doughnuts
Foaming
Forms stable film to entrap gas
Chiffon desserts, cakes, whipped toppings
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Separation technologies in dairy and egg processing 361 whey protein mixtures for the isolation of one or a group of proteins is useful. Considerable progress has been made over the past twenty years in technologies aimed at separation, fractionation and isolation in a purified form of many interesting proteins occurring in both bovine colostrum and milk. Industrial-scale methods have been developed for some proteins but their large-scale manufacture is still limited. Chapter 16 gives more details on fractionation of individual proteins. 12.7.1 Fractionation of individual caseins There is a considerable interest in fractionating whole casein into individual caseins. Potential uses include bovine milk-based infant formulas and preparation of biologically active peptides and specific additives. The native casein micelles retentate obtained from skimmed milk 0.1 mm MF constitutes an excellent raw material for preparing individual caseins. Most studies on the fractionation of the whole casein are related to the isolation of b-casein, the main component of human casein, which contains numerous peptide sequences with high physiological properties. The isolation method is based on the preferential solubilization of the very hydrophobic b-casein at low temperature. At ~4 °C, b-casein dissociates from the casein micelle and can be isolated from the rest of the caseins (caseinate and renneted skimmed milk) using separation techniques such as UF and MF or centrifugation. The yield of b-casein is enhanced at low pH (4.2–4.6) and the b-casein purity can reach 90% (Le Magnen and Maugas, 1991). The co-product (retentate or sediment fractions) is enriched in as and k-caseins. Based on the same principle, a promising process to separate b-casein directly from whole milk has recently been proposed (Lucey and Smith, 2009). It operates in two successive MF steps both using polymeric spiral wound membranes. The first, operating at low temperature, separates b-casein from the retentate containing casein micelles and fat. b-casein is then separated from the rest of the materials present in permeate by warming the permeate to room temperature resulting in aggregation of b-casein. The second MF retain the aggregated b-casein while native soluble proteins passed into the membrane. 12.7.2 Fractionation of serum proteins Several methods related to the isolation of serum proteins have already been published, but many of them are only available at the laboratory scale (Bonnaillie and Tomasula, 2008). Some procedures have, however, led to industrial manufacture of enriched protein fractions or separate proteins such a-lactalbumin, b-lactoglobulin and minor proteins such as LF, LP and immunoglobulins. The cost of extracting these proteins is high but often justified when recognizing the great value-added benefits when incorporated in hygiene products, functional foods and nutraceutical products. Each of the extracted proteins or group of proteins has been proven or implied to have unique
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362 Separation, extraction and concentration processes functional, nutritional or nutraceutical properties. Some putative activities are digestive function (b-lactoglobulin), anticarcinogenic (a-lactalbumin), antimicrobial (LF and LP) and passive immunity (immunoglobulins). Among the commercially interesting proteins, the two main serum proteins, a-lactalbumin and b-lactoglobulin, can be produced in enriched fractions using membrane or centrifugation processes: a-lactalbumin has a great potential market because of its high content of tryptophan (four residues per mole) and in infant milk formula. The main uses of b-lactoglobulin appear to be in gel and foam-type products and in the manufacture of protein hydrolysates for food ingredients. Heat combined with pH adjustment can be exploited to fractionate b-lactoglobulin and a-lactalbumin from whey. a-Lactalbumin is a calcium metalloprotein (1 mol of calcium per mol of protein), that loses its bounded calcium and its stability when at ~55 °C (30 min) the pH is lowered to 3.8. At this pH calcium dissolves in the solution and a-lactalbumin unfolds and precipitates at 50–65 °C. Such physicochemical conditions involve the reversible polymerization of the protein that precipitates together with immunoglobulins and bovine serum albumin (Bramaud et al., 1997). Separation of the precipitate can be performed either by desludging clarifier or microfiltration. The supernatant, or permeate contains the b-lactoglobulin fraction that can be further processed by UF in combination with diafiltration to yield purified b-lactoglobulin (95% purity). The a-lactalbumin (60% purity) can be recovered from the sediment/retentate after solubilization at neutral pH, followed by UF. Starting from the permeate of milk microfiltration, this principle can be used to produce high purity non-lactosylated b-lactoglobulin (Maubois et al., 2001). Ion-exchange chromatography is also used industrially to produce a-lactalbumin and b-lactoglobulin enriched or purified (purity > 90%) fractions (Etzel, 1999; Etzel et al., 2006; Outinen et al., 1996). The high cost of purified fractions often prevents them from being used in targeted food applications. This process features the use of a resin to isolate fraction of protein from the rest of the whey. With a careful choice of the resin system and the eluants a-lactalbumin and b-LG can be separated from each other very precisely. For example, Thuran and Etzel (2004) modified a cationexchange chromatography method previously designed to fractionate sweet whey, to efficiently fractionate acid whey. They used different inexpensive food-grade buffers and the extraction was performed in two steps leading to the release of a-lactalbumin fraction with high recovery rate (96%) and purity (93%) on the one hand and WPI containing mainly b-lactoglobulin on the other hand. Recently, new technologies were shown to produce enriched soluble protein fractions (Bonnaillie and Tomasula, 2009). Supercritical CO 2 fractionation technology is one of the methods proposed for the recovery of large-scale serum protein fractions enriched with both a-lactalbumin and b-lactoglobulin from either WPC or WPI. When supercritical CO2 is injected © Woodhead Publishing Limited, 2010
Separation technologies in dairy and egg processing 363 into solutions containing soluble protein isolate or serum protein concentrate, fractions containing 70% w/w of a-lactalbumin in solid form and 95% w/w of b-lactoglobulin in a soluble liquid form (uncontaminated with chemical additives) were obtained. LF and LP are the two serum proteins having a high isoelectric point (Table 12.2) (Bonnaillie and Tomasula, 2008). Fat normally causes problems for chromatographic separations, as it blocks packed columns as soon as the feed is introduced. Therefore, fat is removed before cation-exchange capture of LF and LP from skimmed milk. They are then classically isolated by ionexchange chromatography from skimmed milk or whey. At the pH of milk or whey, both proteins are specifically adsorbed on cation exchangers, the other proteins being negatively charged. Their elution is generally realized through the use of an increasing ionic strength gradient. Recently, Andersson and Mattiasson (2006) extracted pure LF from WPC using a simulated moving bed chromatographic technology, with several advantages compared with non-moving bed columns (increase in productivity by 48%, increase in LF concentration by 6.5 times, reduction in buffer consumption by 4.3 times). The growing interest in LF is related to its antibacterial properties, by forming an iron complex and inhibiting the growth of micro-organisms by depriving bacteria of iron that is essential for their growth, and to its interesting nutritional activities owing to the transport of iron in the organism. Immunoglobulins can be isolated from whey using an UF membrane with a cut-off about 100 kDa or more but whey is a poor source of immunoglobulins compared with colostrum or milk produced by hyperimmunized cows (Table 12.2). Cow’s colostrum contains substantially higher concentrations of immunoglobulins than mature milk (20–200 g L–1 against 0.15–0.8 g L–1) and then can be used as a appropriate starting fluid for a two-step immunoglobulin extraction procedure (Piot et al., 2004): the colostrum is first microfiltered using a 0.1 mm pore membrane so as to obtain a permeate (named ‘serocolostrum’) which is crystal clear, free of blood and somatic cells as well as fat globules and casein micelles. The permeate that contains 80% of the initial immunoglobulins can then be further concentrated using ultrafiltration (100 kDa). Commercial immunoglobulin products are mostly used in veterinary medicine on neonatal ruminants and pigs. Because ruminants are born without blood antibodies, they are very susceptible to infection and it is highly desirable that they receive protection either by suckling colostrum for at least one week or by ingesting an immunoglobulin concentrate. 12.7.3 Fractionation of peptides Milk is a rich source of bioactive peptides and there is considerable commercial interest in producing bioactive peptides for use in food applications. Bioactive peptides are specific protein fragments, which have a positive impact on body functions and conditions and may ultimately influence health (regulation of weight; mood, memory and stress control; immune defence; and improvement
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364 Separation, extraction and concentration processes of heart, bone, dental and digestive health) (Korhonen, 2009) (Fig. 12.7). The market for bioactive peptides is increasing because the possibility of designing new dairy products with health-promoting benefits looks promising and offers a perspective for consumers and producers. Over the past few decades, a number of methods were developed for their purification (chapter 16) and, to-date, some casein-derived peptides have been manufactured on the industrial scale (Table 12.7). The most common way to produce bioactive peptides is through enzymatic hydrolysis of precursor proteins, using gastrointestinal enzymes, usually pepsin and trypsin. After hydrolysis, the peptides are fractionated and enriched using various methods (precipitation with salts or solvents, ultrafiltration or chromatography). Angiotensin-converting enzyme (ACE)
Egg white pH ª 8.5
Extraction of lysozyme (cation exchange chromatography)
Egg white free of lysozyme
Three steps: • fixation • rinsing (water) • elution (NaCl)
Extracted lysosyme
Crystallization of lysozyme (pH ª 10)
Concentration (filter press)
Solubilization (pH ª 3.5–5.0)
Clarification, concentration, drying (centrifugation, ultrafiltration, filter press, spray drying)
Lysozyme powder
Fig. 12.7 Lysozyme extraction using ion exchange chromatography (adapted from Guérin-Dubiard and Anton, 2010).
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Separation technologies in dairy and egg processing 365 Table 12.7 Marketed bioactive peptides (from Korhonen, 2009) Brand name
Type of product
Claimed functional bioactive peptides
Health/function Manufacturer claims
Calpis Sour milk Val-Pro-Pro, Ile-Pro- Reduction Pro, derived from of blood b-casein and k-casein pressure Evolus Calcium- Val-Pro-Pro, Reduction of enriched Ile-Pro-Pro, derived blood pressure fermented from b-casein milk drink and k-casein Biozate Hydrolysed b-Lactoglobulin Reduction whey protein fragments of blood isolate pressure BioPURE- Whey protein 106–109 fragment Prevention of GMP isolate of k-casein dental caries, influence the clotting of blood, protection against viruses and bacteria PRODIET Flavoured aS1-casein f(91-100) Reduction of F200/ milk drink, (Tyr-Leu-Gly-Tyr-Leu- stress effect lactium confectionery, Glu-Glu-Leu capsules Leu-Arg) Festivo Fermented aS1-casein f(1–6), No health low-fat (1–7), (1–9) claim hard cheese Cysteine Ingredient/ Milk protein- Aids to raise peptide hydrolysate derived peptide energy level and sleep C12 Ingredient/ Casein-derived peptide Reduction of hydrolysate blood pressure Capolac Ingredient Caseinophosphopeptide Helps mineral absorption PeptoPro Ingredient/ Casein-derived peptide Improves hydrolysate athletic performance and muscle recovery Vivinal Ingredient/ Whey protein derived Aids relaxation alpha hydrolysate and sleep Recaldent Chewing Calcium casein Anticariogenic gum peptone–calcium phosphate
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Calpis Co., Japan Valio Oy, Finland
Davisco, USA
Davisco, USA
Ingredia, France
MTT Agrifood Research, Finland DMV International, the Netherlands DMV International, the Netherlands Arla Foods Ingredients, Sweden DSM Food specialist, the Netherlands
Borculo Domo Ingredients (BDI), the Netherlands Cadbury Adams, USA
366 Separation, extraction and concentration processes inhibitory peptides, known for their antihypertensive property, and calciumbinding phosphopeptides, for example, are commonly produced by trypsin (Korhonen and Pihlanto, 2006). For the preparation of phosphopeptides from casein, Brulé et al. (1981) proposed the use of an ultrafiltration membrane for processing permeate after digestion of caseinate in solution with a proteolytic enzyme. The separation of the phosphopeptides present in the permeate was performed by UF of the peptide solution after addition of a bivalent cation salt (calcium chloride) so as to cause aggregation of phosphopeptides. The non-phosphorylated peptides pass through the membrane and diafiltration against water, used to purify the phosphopeptides in the retentate, results in a preparation which is rich (>90% w/w) in the desired phosphopeptides. The glycomacropeptide, the C-terminal part of the k-casein released in whey by the action of chymosin was shown to be separated from sodium caseinate using UF membrane or centrifugation. This peptide has numerous uses, such as action on satiety and inhibition of Escherichia coli cells adhesion to intestinal walls and, in particular, it contains no phenylalanine, which makes it suitable for use as a nutritional protein supplement for patients suffering from phenylketonuria, who cannot digest protein containing phenylalanine owing to their lack of the appropriate degrading enzyme. A number of dairy starter cultures, as well as proteolytic enzymes isolated from lactic acid bacteria also lead to the formation of bioactive peptides from milk proteins, in particular during the manufacture of fermented dairy products. For example, several studies have demonstrated that Lactobacillus helveticus strains are capable of releasing antihypertensive peptides. A UHT milk fermented by the GG strain of Lb. rhamnosus and subsequently digested with pepsin and trypsin produced hydrolysate fractions that were immunosuppressive. The occurrence of many bioactive peptides in bovine milk is now well established (Korhonen, 2009), but currently the industrial-scale production of such peptides is limited by a lack of suitable separation technologies. Among them, membrane techniques, such as NF or UF, are used industrially to produce ingredients that contain bioactive peptides based on casein or whey protein hydrolysates (Table 12.7) and seem to be the best technology available for the enrichment of bioactive peptides.
12.8 Treatment of effluents and technical fluids in the dairy industry In many countries, the dairy industry is considered to be one of the largest generators of food-processing wastes. Owing to large water volumes for factory cleaning and disinfection, the dairy industry produces 0.2–11.0 L effluents/L of processed milk. Cleaning-in-place (CIP) operations significantly contribute to water consumption and are responsible for 50 to 95% of the overall volume of waste streams (Daufin et al., 2000, Alvarez et al., 2007). © Woodhead Publishing Limited, 2010
Separation technologies in dairy and egg processing 367 The polluting load of these effluents ranges from 0.2 to 5.0 g L–1 of chemical oxygen demand (COD), and is mainly caused by loss of raw material (0.5 to 2% of milk). CIP systems are also mainly responsible for the high pH value (9–11) of the end-of-pipe wastewaters. Although a high portion of dairy wastewaters is still land spread, the treatment of the effluents mainly takes place in biological treatment plants at most dairies. For a factory of average capacity (106 L of milk a day), the sludge produced is usually used for land spreading (1 to 3 tonne of dry matter) and purified water drained to rivers (0.3 ¥ 106 to 3 ¥ 106 L). Current European regulations relating to landfill management, land spreading and purified water quality along with social pressure, has recently forced the dairy industry to significantly reduce its production of sludge and to improve purified water quality. Over the past few years, the dairy industry has been attempting to find new concentration and separation processes to reduce its effluents production. Several types of effluents are currently treated in the industry, benefiting mainly from the potential of membrane technology: washing waters of rennet casein precipitate were treated, respectively, using NF and dead-end filtration; white (flushing) and pre-rinsing waters, corresponding to the first step of CIP, were treated using UF, NF or RO (Blanchard, 1991; Delbecke, 1981). The outcome is a highly significant improvement of water quality after treatment [suspended solids (SS) < 2 mg L–1; chemical oxygen demand (COD) <35 mg L–1; BOD5 <3 mg L–1) and the re-use of milk components (either as recycling back to the production unit or animal feed). Evaporation condensates (also called ‘cow’s water’) and permeates of milk and whey NF with COD of 10–1000 mg L–1, can be treated by RO with eventual UF or MF pretreatment (Horton, 1997). In both cases, the RO permeate produced can be used as a source of water with ‘food quality’, for rinsing and cleaning operations. Cheese brines (170–230 g kg–1 NaCl) are widely recycled after a UF (50 kDa) but more commonly after a MF (0.2–1.4 mm) step, which strongly reduce microbiological counts, without altering the chemical composition in contrast to conventional pasteurization (Pedersen, 1992). Heat treatment and Kieselguhr (diatomaceous earth) filtration are still the most well-used technologies for brine treatment but the former modifies the calcium phosphate equilibrium in solution, and the latter is recognized by The World Health Organization as a cause of lung disease, thus requiring safe working conditions. Finally, used alkaline and acid CIP solutions, which are periodically drained to waste, can be advantageously regenerated. The automatic renewal of cleaning solutions once a week, which is generally practised, leads to an excessive consumption of water and chemicals, and mainly caustic soda (about 120 tonne/year for a plant producing 106 L day–1 of milk). Desludging of caustic soda solutions is currently practised, but MF, which retains suspended solids and transmits small molecules responsible for the low surface tension of the re-used solutions, seems to be an appropriate operation both from the
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368 Separation, extraction and concentration processes technical and the economical points of view (Gésan-Guiziou et al., 2007). From the beginning of the cleaning process and regardless of the equipment used and the type of CIP processing, low values of surface tension (attributed to chemical reactions of proteins and fat in alkaline conditions) were observed in caustic soda solutions when their loading charge was increased (Alvarez et al., 2007). Because the low surface tension characteristics led to high cleaning kinetics (Alvarez et al., 2007), the MF-regenerated caustic soda solutions is as fast as that of a commercial alkaline detergent, and much more efficient than fresh caustic soda.
12.9 Conclusions and future trends in the dairy industry Membrane separation technologies and, to a lesser extent, chromatography, offer to the dairy technologist several techniques for the extraction and purification of almost all the main proteins of milk and for the treatment of wastewaters. Membrane operations in particular have made it possible for new and original products to be created. The fractionation of casein micelles from serum proteins using MF membrane is one of the best examples of such a success, which would have been considered unlikely some years ago. The differences in attributes between serum proteins obtained from milk (using milk MF) and those isolated from cheese whey are becoming an important factor for dairy protein processors considering, very recently, this new avenue for fractionating proteins and for producing bioactive peptides. Apart from being a balanced source of valuable amino acids, milk proteins contribute to the specific properties of various dairy products, and possess interesting biological properties, which make them potential ingredients of health-promoting foods. A few commercial protein and peptide fractions have been launched on the market and this trend is likely to continue alongside increasing knowledge about the functionalities of the products. Advances in membrane design, the best understanding of the limiting phenomena occurring during filtration operations, and the recent decrease of the processing cost of polymer membrane, make this technology more and more attractive. Therefore further integration of membrane operations is to be expected, provided they are designed in such a way that at each processing step, membrane fouling is limited, membrane cleaning is optimized, and end products, co-products and wastes are given equal attention. Through the dairy industry, membrane processes have been shown to provide the food industry with efficient tools for limiting the environmental impact of the food sector. There is no doubt that in the near future any food process will include at least one membrane operation for their effluents treatments. New applications are also likely to be developed, such as recovery of phospholids derived from the fat-globule membranes from buttermilk (aqueous phase produced from churning butter) and recovery of growth factors present in cows’ colostrum. Moreover, new emerging technologies, © Woodhead Publishing Limited, 2010
Separation technologies in dairy and egg processing 369 such as rotating and vibrating filtration, emulsification, and supercritical carbon dioxide fractionation, have met with some success in the laboratory and this could lead to more commercial applications in the future.
12.10 The egg products industry and composition of egg products Owing to their high functionalities (foaming, coagulative, emulsification, and binding), many egg products are used as ingredients in various food applications, such as bakery products, meringues, mayonnaise, cookies and meat products. Egg white in particular is in great demand in the baking, confectionery and cake mix industries for its whipping and foaming properties and it is mainly used in the form of dried egg white solids. Moreover, eggs are an excellent source of high-quality proteins containing all nine essential amino-acids and other nutrients required in our diets, and possessing bioactive properties including antimicrobial activity, protease inhibitory action, immunomodulatory, anticancer and antihypertensive activities, vitamin binding properties and antigenic or immunologic characteristics. Knowledge of the characteristics of the individual proteins has led to the development of new separation processes now industrially used for the extraction of three proteins of egg white, and of some components of the yolk. Eggs are composed of three main parts (Burley and Vadehra, 1989): the eggshell with the eggshell membrane, the albumen also named egg white, and the yolk. The yolk is surrounded by the white, which in turn is enveloped by eggshell membranes and finally by a hard eggshell. The egg white makes up about 66% of the liquid weight of the egg. It contains about 88–90% water. Proteins are the major components of albumen solids (about 10–11% of the white weight) whereas carbohydrates (mostly free glucose) (≈ 0.8–1.0%), minerals (≈0.5%) and lipids (≈0.03%) are minor components. Table 12.8 lists the content and some characteristics reported for some proteins in egg white. It should be noticed that despite the numerous recent studies for separating and identifying the proteins located in hen’s egg, many proteins remain uncharacterized or even unknown. Except for lysozyme and avidin, most proteins are acid proteins and are negatively charged at the natural pH of the egg white (pH 9.0–9.3). Glucose is the main ‘free’ sugar, and is usually removed by bacterial fermentation and enzyme hydrolysis before drying of egg white. In some applications, the presence of glucose is undesirable because it causes a detrimental effect on storage stability and quality of the product, by the forming of off-flavours and brown pigments caused by Maillard reaction. Egg yolk contains about 48–51% water, according to the age of the laying hen and the duration of preservation. Lipids are the main components (31% of the total weight) of the egg yolk solids. The lipid distribution is 65% © Woodhead Publishing Limited, 2010
370 Separation, extraction and concentration processes Table 12.8 Content and characteristics (isoelectric point, molecular weight) of some proteins found in egg white (from Li-Chan and Kim, 2008) Protein
Protein % of Isoelectric point egg white
Ovalbumin 54 Ovalbumin Y Ovotransferrin 12 Ovomucoid 11 Ovomucin 3.5 Lysozyme 3.4 Ovoglobulin G2 globulin 4.0 G3 globulin 4.0 Ovoinhibitor 1.5 Ovoglycoprotein 1.0 Ovoflavoprotein 0.8 Ovomacroglobulin 0.5 Cystatin 0.05 Avidin 0.05
4.5 (5.1–5.3) (5.3–5.5) 6.1 (6.2–6.7) 4.1 4.5–5.0 10.7 (6.1–5.3) 5.5 4.8 5.1 (6.2–6.4) 3.9 (5.0–5.4) 4 (5.0–5.2) 4.5 5.1 (6.1) 10
Molecular weight (kg mol–1 or kDa) 45 (42.4) (53.4–54.3) 76 (85–75) 28 (37.2–43.1) 5500–8300 14.3 (15) 30–45 Not determined 49 (69.5–63.6) 24.4 (37.2–43.1) 32 (37.4–40) 769 12.7 (17) 68.3
Source: data were compiled from Li-Chan et al. (1995), except for those in parentheses, which are from Guérin-Dubiard et al. (2006).
triglycerides, 28–30% phospholipids, and 4–5% cholesterol. Yolk contains about 16% proteins, consisting of livetins (globular proteins) and lipoproteins particles including low and high-density lipoproteins. As a food, yolk is a major source of vitamins and minerals (3.5% of dry yolk). Whole egg contains about 25% solids, 23% proteins and 10% fat. Minor amounts of minerals and carbohydrate are also present.
12.11 Concentration and stabilization of egg white and whole egg The concentration of egg white and whole egg (it is not necessary to concentrate egg yolk, which contains about 50% solids) is used in two main industrial applications: the concentration of egg components before spray drying and the stabilization of egg white and whole egg. Owing to its high water content, the egg white is quasi-systematically concentrated before spray drying to reduce energy costs. Egg products are rarely concentrated using conventional thermal evaporation because of the damage caused to the highly heat- and shear-sensitive egg white proteins. Membrane processes, especially RO and UF with low molecular weight cutoff (MWCO) in order to limit the loss of nitrogen matter, were investigated as a means of concentrating egg before drying (Bergquist, 1995). These processes have been commercially applied to the concentration of egg white.
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Separation technologies in dairy and egg processing 371 RO concentrated all egg components and led to a permeate, free of organic matter, that could be re-used for intermediate cleaning of operations. UF is sometimes more useful because it not only removed water, but also lowered glucose, sodium and potassium levels by about 50% in the concentrated product. The partial removal of glucose reduced problems observed during storage and excessive browning during baking processes. A higher removal of glucose, while retaining the egg white protein, can be achieved by carrying out UF in conjunction with diafiltration. UF led also to higher permeation fluxes, lower energy consumption (owing to lower applied TMP), superior functional properties of the concentrates owing to the removal of free glucose and salts. Neither of the concentration methods (RO or UF) significantly affected foaming properties, and UF was observed to improve gel strength of egg white, which was probably related to the increase in protein concentration. The concentration of whole egg using UF is marginal. The concentration of liquid egg white and whole egg is used for their stabilization. The objective of this operation is to concentrate the egg product by UF and then add salt and/or sugar to decrease the water activity (aw) of the product under the thresholds of micro-organisms development (aw = 0.85) (Bonduelle, 1978; Liot 1980). This process enables the processing of concentrated egg white up to 33% dry matter content with a maximum sugar content of 50% or salt content of 9% (0.80
12.12 Industrial extraction of egg-white proteins Owing to the high nutritional, functional and biological properties of eggwhite proteins, many studies are currently in progress for the extraction of
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372 Separation, extraction and concentration processes albumen proteins. Several methods of proteins isolation have already been proposed, but very few have been developed on an industrial scale (Nau et al., 2010). Lysozyme, ovotransferrin and avidin are the three proteins commercially removed from egg white. 12.12.1 Lysozyme Lysozyme is the only hen’s egg protein routinely used commercially. It can be used as a preservative and finds practical applications in the food, pharmaceutical and medicine sectors (Lesnierowski and Kijowski, 2007). For example this protein is used in cheese to prevent contamination because it does not inhibit the starter and secondary cultures required for the ripening of the cheeses. It demonstrates antibacterial properties, particularly against Gram-negative bacteria, among them a number of food pathogens such as Listeria monocytogenes. Lysozyme is known as a hydrolysate that cuts the b-1-4 linkage of the glycosidic bond between polysaccharide copolymers, which represent structural units of many bacterial cell walls. Lysozyme from hen’s egg white is a small protein with a molecular weight of 14.2 kDa having, apart from the other egg albumen proteins, a strong basic character (isoelectric point 10–11, Table 12.8). Both physical and chemical properties of lysozyme have been exploited as isolation processes. The small size of the protein was used in membrane techniques, particularly UF or MF, to separate lysozyme from egg white. However, the ability of this enzyme to electrostatically bind with other negatively charged egg white proteins greatly reduced the transmission of the protein through the membrane. Despite these membrane processes being tested on a large scale (Peri and Feriscini, 1972, Lepienne et al., 1986, Kijowki et al., 1998), they are not yet developed on an industrial scale. Currently, the commercial procedures of lysozyme isolation focus on two main processes, based on the very high isoelectric point of lysozyme: the selective precipitation of lysozyme and extraction using ion-exchange chromatography. The selective precipitation of lysozyme combines adjustment of the pH to 10 (leading to a pH close to the isoelectric point of the protein) and addition of 5% sodium chloride to increase ionic strength and favour precipitation. With this technique, several precipitations and resolubilizations are necessary to obtain high-purity protein. The remaining egg white after lysozyme separation becomes difficult to valorize owing to the high residual content of salt. It is possible to reduce its salt content by using UF and diafiltration techniques and desalted remaining egg white has been shown to maintain foaming properties when compared with the native egg white. Ion-exchange chromatography and, in particular, cation-exchange chromatography has mainly been used on an industrial scale to isolate lysozyme. In association with salt precipitation of the extracted lysozyme, this technique can lead to a very pure protein fraction (>99%) and to a co-
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Separation technologies in dairy and egg processing 373 product with high techno-functional properties. Figure 12.7 gives the main steps of the industrial process: after slight acidification of the egg white down to 8.5, the egg white was put into contact with ion-exchange resins, such as Amberlite, carboxymethylcellulose (CMC), carboxymethyl-Sephadex, or Duolite. Both column and batch techniques are used for large-scale operations. In the batch system, resin is stirred so as to maintain it in suspension and favour the sorption of lysozyme from egg white. When the lysozyme is fixed on the resin, the egg white free of lysozyme is extracted; the resin is washed with water before the elution of the protein with sodium chloride solution. Crystallization of lysozyme, initiated with already formed crystals, is then performed at a pH close to 10 in order to increase its final purity. The final lysozyme crystals are then concentrated by press filter, solubilized at acid pH (3.5–5.0), clarified by centrifuge, concentrated by UF, filtered again with a press filter for bacterial removal and finally dried. 12.12.2 Ovotransferrin Ovotransferrin is produced on an industrial scale, but to a lesser extent than lysozyme (Guérin-Dubiard and Anton, 2010). This protein is similar to serum transferrin in animals having the functions and interesting nutritional activities as a result of the transport of iron in organisms. Many procedures, mainly based on liquid chromatography as for all proteins belonging to the transferrin family, have been developed to purify this protein. Owing to its neutral isoelectric point (pI 6.2–6.7, Table 12.8) the extraction can be performed either by cation or by anion-exchange chromatography. When performing cation-exchange chromatography, lysozyme should be first extracted. When performing anion-exchange chromatography, all the proteins with an isolectric point lower than the ovotransferrin one should be first separated. Yields of extraction vary from 50 to 80%, and ovotransferrin purity can reach 98% (Guérin-Dubiard and Anton, 2010). One can note that the extraction of ovotransferrin can be performed from egg-white free of ovomucin which can easily be obtained after centrifugation of the egg-white after dilution with water (1:4) and adjustment of the pH at 6. Ovomucin has the ability to precipitate at low ionic strength, which leads to serious problems during water washing phases of the chromatography process. The extraction of ovotransferrin from the egg-white free of ovomucin can lead to a protein purity up to 90%. 12.12.3 Avidin Avidin represents a maximum of 0.05% of the total protein content of eggwhite (Table 12.8). It is best known for its binding properties with biotin (vitamin H or B8), an essential growth factor. Owing to its high biotin affinity, avidin is thought to serve as a defensive protein against biotinrequiring micro-organisms. It is also widely used as tool in a number of
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374 Separation, extraction and concentration processes affinity-based separations, in biochemical diagnostic assays, and in a variety of other applications. Several isolation processes based on chromatographic separations have been proposed either based on ion-exchange or affinity properties. Avidin is an alkaline glycoprotein similar to lysozyme (Table 12.8), and then its isolation can be performed using processes based on ion- (mainly cation-) exchange chromatography. However, the extracted avidin is often contaminated by lysozyme. Several authors proposed adapted procedures in order to increase the selective elution of avidin fixed on the cationic support (Nau et al., 2007). Durance and co-workers proposed a novel ion-exchange chromatography procedure depending on pH, ionic strength and/or nature of the salts used in the eluting buffer (Durance and Nakai, 1988; Durance et al., 1991). This procedure made it possible to simultaneously extract avidin and lysozyme from undiluted egg white. Avidin recovery was around 75% and the purity rate of the avidin could reach 40%. The purity of avidin was increased by a subsequent chromatography step. Rao et al. (2003) proposed another original procedure: the selective elution of avidin fixed on a Streamline SP (sulfopropyl) support was performed using hydroxyazobenzene-2’-carboxylic acid (HABA). Avidin, which possesses a high affinity for this molecule, is the only protein being eluted, leading to a very high purity (98%). Owing to its high biotin affinity, avidin is also purified using affinity chromatography. The high commercial value of this protein makes it acceptable to extract this protein with the costly but very selective affinity processes. The recovery and purity of the protein obtained are very high (> 95%), but the materials are very expensive and their lifetimes are rather limited, which could explain why affinity chromatography is classically not used for large-scale separations. Garret-Flaudy and Freitag (2000/2001) proposed an affinity precipitation process using iminobiotin–polymer, which gave an avidin purity higher than 90%.
12.13 Industrial extraction of yolk components The fractionation of yolk components does not exclusively concern the isolation of proteins, as previously seen for the egg white. The g-livetin (egg antibody which corresponds to the immunoglubulin of yolk, immunoglubin Y, IgY), and the phospholipids mainly composed of phosphatidylcholine, PC (80–85% commonly named ‘lecithin’), and phosphatidylethanolamine, PE (10–15%), are commercially extracted. g-Livetin has been exploited on an industrial scale for nearly twenty years because its extraction is easier than immunoglobulins from mammalian blood. This egg antibody is promising for use in immunoassays, to quantify toxins or pathogenic viruses for example, and its use as a functional tool in pharmaceutical applications is developing. Various extraction methods were reviewed in detail by Schade et al. (2005). Two main isolation procedures © Woodhead Publishing Limited, 2010
Separation technologies in dairy and egg processing 375 have been developed. The first method comprises the precipitation of the g-livetin by polyethyleneglycol and the second one uses ion-exchange chromatography leading to 70% recovery with 60% purity. In both cases, the protein is first recovered in the soluble phase of the yolk by water dilution (6 times; pH 5, 4 °C for 6 h). Egg-yolk phospholipids offer several industrial applications, mainly in the nutritional, pharmaceutical and cosmetic fields. This fraction is ultimately used in the food industry as an emulsifier, viscosity reducer, lubricant, and as an antispattering, wetting and release agent, and its widespread uses are linked to its ability to act as a surface-active substance in multiphase systems. Phospholipids from yolk could also serve as encapsulation systems (liposomes and double emulsions) intended for cosmetic formulas and medical uses. They have several other properties including antioxidative activity and inhibition of cholesterol absorption. Industrial extraction methods for phospholipids are commonly based on organic solvents. Phospholipids are soluble in hydrocarbons and other organic solvents, whereas they are typically insoluble in acetone. The latter characteristic allows the separation of the accompanying lipids in order to increase the purity of the phospholipids. Ethanol is also used to increase PC and PE because they are highly soluble in it. On this basis, Juneja et al. (1994) proposed a method for large-scale preparation of phospholipids, treating fresh egg yolk by a combination of acetone and ethanol additions and filter press separations. This process, which produced phopholipid fraction containing the same proportion of individual phospholipids as in the initial yolk (80–85% PC; 10–15% PE), can be extrapolated to the industrial scale (Guérin-Dubiart and Anton, 2010; Juneja et al., 1994). However, because the use of organic solvents in common extraction procedures is questionable on the grounds of safety, alternative techniques for example using supercritical carbon dioxide fluid have been studied and gave positive results (Sim, 1994).
12.14 Conclusions and future trends in the egg-processing industry The egg-processing industry has made a great deal of progress since the 1950s, but current research by food scientists is still adding to the understanding of the components of eggs and their fractionation. The recent isolations on a laboratory scale and characterizations of egg components give clear evidence of the potential uses of eggs: the protein and lipid fractions of egg white and yolk not only provide interesting functional properties but also offer real opportunities for nutraceutical and human health applications. Efforts are still required to identify and produce molecules of interest for exploiting the potential of egg-derived molecules.
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376 Separation, extraction and concentration processes The egg-processing industry has recently learned to fractionate their products. Currently lysozyme is the only egg component routinely removed from egg. The separations used to produce ovotransferrin, avidin, IgY and phospholipids enriched fractions are likely to be developed, but more potential utilization of egg can be envisaged, mainly in biotechnology processes (chapter 19). Enzymatic hydrolysis of egg white to produce various peptides is promising. Peptides have bacteriostatic or antioxidant activities; some lower blood pressure of hypertensive rats. Thus the egg-processing industry has a strong base to expand into new innovative applications, with particular emphasis on the heath benefits of eggs. Finally, for environmental reasons, the treatment of the wastewater produced by the egg-processing industry is likely to develop in the next few years. The effluents issued from the eggbreaking machines for instance is particularly high with a BOD5 ranging from 1–22 g L–1. The membrane technologies will certainly offer good opportunities to decrease the load of effluents sent to the purification plant, in a similar way to that previously done in the dairy industry.
12.15 Sources of further information and advice Two clear and comprehensive books, which emphasize basic aspects of all kinds of membrane processes, but do not give much detail on dairy and egg products applications: ∑ ∑
Zeman LJ and Zydney AL (1996), Microfiltration and ultrafiltration: principles and applications, New York, Marcel Dekker; Cheryan M (1998), Ultrafiltration and microfiltration handbook, Lancaster, Technomic Publishing.
More detailed information about principles of cheesemaking and general aspects of dairy science in: ∑
Walstra P, Wouters JTM and Geurts TJ (2006), Dairy science and technology, Boca Raton, FL, CRC Press, Taylor and Francis.
More detailed information of membrane processes applied to dairy fluids in: ∑
Britz TJ and Robinson RK (2008), Advanced dairy science and technology, Oxford, Blackwell Publishing; ∑ Mistry VV and Maubois JL (2004), ‘Application of membrane separation technology to cheese production’, in Fox PF, McSweeney PLH, Cogan TM and Guinee TP, Cheese chemistry, physics and microbiology. Vol 1 General aspects, London, Elsevier.
More detailed information about egg composition and functionalities in: ∑
Huopalahti R, Lopez-Fandino R, Anton M and Schade R (2007), Bioactive egg compounds, Berlin, Springer; © Woodhead Publishing Limited, 2010
Separation technologies in dairy and egg processing 377 ∑
Mine Y (2008), Egg bioscience and biotechnology, Hoboken, John Wiley & Sons Inc.
More detailed information about egg processing in: ∑
Nau F, Guérin-Dubiard C, Baron F and Thapon JL (2010), Science et technologie de l’œuf et des ovoproduits. Vol 1: Production et qualité de l’œuf; Vol 2: De l’œuf aux ovoproduits, Paris, Tec&Doc Lavoisier (in press).
12.16 References Ahmad S, Gaucher I, Rousseau F, Beaucher E, Piot M, Grongnet JF and Gaucheron F (2008), ‘Effects of acidification on physico-chemical characteristics of buffalo milk: a comparison with cow’s milk’, Food Chem, 106, 11–17. Alvarez N, Gésan-Guiziou G, Daufin G (2007), ‘The role of surface tension of reused NaOH on the cleaning efficiency in dairy plants’, Int Dairy J, 17, 404–411. Andersson J, Mattiasson B (2006), ‘Simulated moving bed technology with a simplified approach for protein purification – separation of lactoperoxidase and lactoferrin from whey protein concentrate. J Chromatogr A, 1107(1–2), 88–95. Bergquist DH (1995), ‘Egg dehydration’, in Stadelman WJ and Cotterill OJ, Egg science and technology, Haworth Press, Fourth edition, Binghamton, NY, 335–369. Blanchard BD (1991), ‘Plant effluents dairy waste streams recovery’, Dairy Food Environ Sanit, 11(9), 494–496. Bonnaillie L, Tomasula PM (2008), ‘Whey protein fractionation’, in Onwulata CI and Huth PJ, Whey processing, functionality and health benefits, Wiley Blackwell, Singapore, 15–39. Bonnaillie L, Tomasula PM (2009), ‘Supercritical carbon dioxide fractionation of serum protein isolate for new food-grade ingredients’, IFT ‘09 Annual Meeting & Food Expo. Paper No. 251: 14. Bonduelle M (1978), ‘Une nouvelle technique dans la conservation des ovoproduits’, Ind Agric Aliment, 95, 1043–1048. Bramaud C, Aimar P, Daufin G (1997), ‘Whey protein fractionation: isoelectric precipitation of a-lactalbumin under gentle heat treatment’. Biotechnol Bioeng, 56(4), 391–397. Brulé G, Roger L, Fauquant J, Piot M (1981), ‘Phosphopeptides from casein-based material’, US patent, 4358465. Burley RW, Vadehra DV (1989), The avian egg: Chemistry and biology, New-York Wiley Interscience. Daufin G, Gésan-Guiziou G, Boyaval E, Buffière P, Lafforgue C, Fonade C (2000), ‘Minimization des rejets liquides de l’industrie laitière par traitement des effluents à l’aide de procédés à membrane’, Tribune de l’eau, 600(4), 25–33. Delbecke R (1981), ‘Recovery of milk by hyperfiltration’, Milchwissenschaft, 36(11), 669–672. Durance TD, Nakai S (1988), ‘Simultaneous isolation of avidin produced by binding of biotin’, J Food Sci, 53, 1096–1106. Durance T, Li-Chan E, Nakai S (1991), ‘Process for the isolation and separation of lysozyme and avidin from egg white’, Canadian Patent 1,283,072. Etzel MR (1999), ‘Isolating b-lactoglobulin and a-lactalbumin by eluting from cation exchanger without sodium chloride’, US patent 5,986,063. Etzel MR, Helm TR, Vyas HK (2006), ‘Methods and compositions involving whey protein isolates’, US patent, 60,569,078.
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378 Separation, extraction and concentration processes Foegeding EA, Zulweska DM, Barbano DM, Drake MA, Luck PJ, Yong YH, Vardhanabhuti B, Berry T (2009), ‘Comparison of the functional properties of serum proteins isolated from milk or whey’ J Dairy Sci, 92(E-Suppl. 1), 163. Froning GW (2008), ‘Egg products industry and future perspectives’, in Mine Y Egg bioscience and biotechnology, Wiley Interscience, Hoboken, New Jersey, US. Garret-Flaudy F, Freitag R (2000/2001), ‘Use of avidin-(imino) biotin system as a general approach to affinity precipitation’, Biotechnol Bioeng, 71, 223–234. Gaucheron F (2005), ‘The minerals of milk’, Reprod Nutr Dev, 45, 473–483. Gésan-Guiziou G (2010), ‘Removal of bacteria, spores and somatic cells by centrifugation and microfiltration techniques’, in Griffiths M Improving the safety and quality of milk, Woodhead Publishing Ltd, Cambridge. Gésan-Guiziou G, Alvarez N, Jacob D, Daufin G (2007), ‘Cleaning-in-place coupled with membrane regeneration for re-using caustic soda solutions’ Sep Purif Technol, 54, 329–339. Gésan-Guiziou G, Boyaval E, Daufin G (1999), ‘Critical stability conditions in crossflow microfiltration of skimmed milk: transition to irreversible deposition’, J Membr Sci, 158, 211–222. Goudédranche H, Fauquant J, Maubois JL (2000), ‘Fractionation of globular milk fat by membrane microfiltration’ Lait, 80, 93–98. Guérin-Dubiard C, Anton M (2010), ‘Fractionnement de l’œuf’, in Nau F, Guérin-Dubiard C, Baron F and Thapon JL Science et technologie de l’œuf et des ovoproduits Vol 2: De l’œuf aux ovoproduits, Lavoisier Tec&Doc, Paris. Guérin-Dubiard C, Pasco M, Molle D, Desert C, Croguennec T, Nau F (2006), ‘Proteomic analysis of hen egg white’. J Agric Food Chem, 54(11), 3901–3910. Hofland GW, Berkhoff M, Witkamp GJ, Van der Wielen (2003), ‘Dynamics of precipitation of casein with carbon dioxide’. Int Dairy J, 13(8), 685–697. Horne DS (2006), ‘Casein micelle structure: models and muddles’. Curr Opinion Colloid Interface Sci, 11, 148–153. Horton BS (1997), ‘Water, chemical and brine recycle or reuse – applying membrane processes’ Austr J Dairy Technol, 52(1), 68–70. Juneja LR, Sugino H, Fujiki M, Kim M, Yamamoto T (1994), ‘Preparation of pure phospholipids from egg yolk’, in Sim JS and Nakai E, Egg uses and processing technologies – new developments, CAB International, Wallingford, UK, 139–149. Kijowki J, Lesnierowski G, Fabisk-Kijowska A (1998), ‘Methods of lysozyme separation, enzyme molecular form and functional quality of the residual egg white’, in The 2nd international symp egg nutrition newly emerging ovotechnologies, Banff, Alberta, 54. Korhonen H (2009), ‘Milk-derived bioactive peptides: from science to applications’ J Funct Foods, I, 177–187. Korhonen H, Pihlanto A (2006), ‘Bioactive peptides: production and functionality’ Int Dairy J, 16, 945–960. Korhonen H, Pihlanto A (2007), ‘Technological options for the production of healthpromoting proteins and peptides derived from milk and colostrum’, Curr Pharm Des, 13, 829–843. Le Magnen C, Maugas JJ (EURIAL) (1991), ‘Method and device for obtaining beta casein’, Patent PCT / FR 91/00506. Lepienne A, Maubois JL, Thireau M, Piot M (1986), ‘Procédé pour l’obtention de lysozyme par microfiltration à partir d’une matière à base de blanc d’œuf’, French patent 25697222. Lesnierowski G, Kijowski J (2007), ‘Lysozyme’, in Huopalahti R, Lopez-Fandino R, Anton M and Schade R Bioactive egg compounds, Springer, Berlin, pp. 33–42. Li-Chan ECY, Kim HO (2008), ‘Structure and chemical composition of eggs’, in Mine Y Egg bioscience and biotechnology, Wiley Interscience, Hoboken, New Jersey, USA 1–95.
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Separation technologies in dairy and egg processing 379 Li-Chan ECY, Powrie WD, Nakai S (1995), ‘The chemistry of eggs and egg products’, in StadelmanWJ and Cotteril OJ, Egg science and technology, Haworth Press, Fourth edition, Binghamton, NY, 105–175. Liot R (1980), ‘Produit hautement concentré de blanc d’œuf ou d’œuf entier salé et son procédé de préparation’, French patent 8022309. Lucey J, Smith K (2009), ‘An integrated processing system to produce beta-casein, native serum protein and casein concentrates from whole milk’, J Dairy Sci, 92(ESuppl.1), 164. Marshall KR (1982), ‘Industrial fractionation of milk proteins: serum proteins’, in Fox PF Developments in dairy chemistry-1, Applied Science Publishers, NY. Maubois JL (1981), ‘Perspectives d’utilisation des techniques à membranes dans les industries agro-alimentaires’. Académie d’Agriculture de France, Procès-verbal 26 nov. 1980, 1451–1461. Maubois JL, Fauquant J, Famelart MH, Caussin, F (2001), ‘Milk microfiltrate, a convenient starting material for fractionation of whey proteins and derivates’, in Proceeding of the 3rd International Whey Conference, Munich Sept 12–14, Behr’s Verlag, Hamburg. Maubois JL, Ollivier G (1997), ’Extraction of milk proteins’, in Damodaran S and Paraf A, Foods proteins and their applications, New York, Marcel Dekker Inc, 579–595. Michalski MC, Camier B, Gassi JY, Briard-Bion V, Leconte N, Famelart MH, Lopez C (2007), ‘Functionality of smaller vs control native milk fat globules in Emmental cheeses manufactured with adapted technologies’, Food Res Int, 40, 191–202. Michalski MC, Leconte N, Briard-Bion V, Fauquant J, Maubois JL, Goudédranche H (2006), ‘Microfiltration of raw whole milk to select fractions with different fat globule size distributions: process optimization and analysis’, J Dairy Sci, 89(10), 3778–3790. Mistry VV, Maubois JL (2004), ‘Application of membrane separation technology to cheese production’, in Fox PF, McSweeney PLH, Cogan TM and Guinee TP, Cheese chemistry, physics and microbiology. Vol 1 General aspects’ London, Elsevier. Mulder H, Walstra P (1974), ‘The milk fat globule. Emulsion science as applied to milk products and comparable foods’, Commonwealth Agricultural Bureaux, Farnham Royal, UK. Nau F, Guérin-Dubiard C, Baron F, Thapon JL (2010), ‘ Science et technologie de l’œuf et des ovoproduits’ Vol 2: De l’œuf aux ovoproduits, Paris, Tec&Doc Lavoisier (in press). Nau F, Guérin-Dubiard C, Croguennec T (2007), ‘ Avidin’, in Huopalahti R, Lopez-Fandino R, Anton M and Schade R Bioactive egg compounds, Springer, Berlin, 75–80. Noël R (Société Vidaubanaise d’Ingénierie) (1992), ‘Procédé de séparation du phosphocaséinate de calcium et du lactosérum d’un lait écrémé plus généralement d’un composé protéique d’un liquide biologique’. PCT WO 92/12642. Outinen M, Tossavainen O, Syväoja EL (1996), ‘Chromatographic fractionation of a-lactalbumin and b-lactoglobulin with polystyrenic strongly basic anion exchange resins’, Lebens Wis Technol, 29(4), 340–343. Pearce RJ (1992), ‘Protein recovery and whey protein fractionation’. in Zadow JG Whey and lactose processing, Elsevier Applied Science, 271–316. Pedersen PJ (1992), ‘Microfiltration for the reduction of bacteria in milk and brine’ In: ‘New applications of membrane processes’, Int Dairy Fed Bulletin, Special issue 9201, 33–50. Peri C, Feriscini C (1972), ‘Concentration and fractionation of egg white by ultrafiltration’ Sci Technol Alimenti, 2, 120–122. Piot M, Fauquant J, Madec MN, Maubois JL (2004), ‘Preparation of “serocolostrum” by membrane microfiltration’, Lait, 84, 333–342. Qi P X (2007), ‘Studies of casein micelle structure: the past and the present’, Lait, 87, 363–383. Quiblier JP, Ferron-Baumy C, Garric G, Maubois JL (1991), ‘Procédé de traitement des
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380 Separation, extraction and concentration processes laits permettant au moins de conserver leur aptitude fromagère’, Patent FR 2 681 218 A1. Rao M, Gupta M, Roy I (2003), ‘Process for the isolation and purification of a glycoprotein avidin.’ International patent 03/099035 A1. Sandblom RM (Alfa-Laval) (1974), ‘Filtering process’, Swedish patent 7,416,257. Schade M, Calzado EG, Sarmiento R, Chacana PA, Porankiewicz-Asplund J, Terzolo HR (2005), ‘Chicken egg yolk antibodies (IgY-technology): a review of progress in production and use in research and human and veterinary medicine’. ATLA, 33, 129–154. Sim JS (1994), ‘New extraction and fractionation method for lecithin and neutral oil from egg yolk’, in Sim JS, and Nakai E, Egg uses and processing technologies – new developments CAB International, Wallingford (UK), 128–138. Stadelman WJ, Cotteril OJ (1995), Egg science and technology, 4th ed. Binghamton, NY: Haworth Press. Tamine AY, Robinson RK (2007), ‘Yoghurt: science and technology’. CRC Press, Abington, England. Thuran KN, Etzel MR (2004), ‘Whey protein isolate and a-lactalbumin recovery from lactic acid whey using cation-exchange chromatography’, J Food Sci, 69(2), 66–70.
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Separation technologies in the processing of fruit juices 381
13 Separation technologies in the processing of fruit juices G. Vatai, Corvinus University of Budapest, Hungary
Abstract: After an introduction to the fluids in the fruit juice product sector, the various separation processes for production of fruit juices from different fruits are discussed. The separation processes for the production of fruit juice concentrate are reviewed, and the advantages and disadvantages of the applied techniques are discussed. Finally, a multistep membrane process for must concentration is described using laboratory experimental data, modeling and optimization. Key words: fruit juice, separation, extraction, membrane filtration, membrane distillation, osmotic distillation.
13.1 Introduction With the changing of nutrition trends, interest has focused on plants and crops with many valuable components. Fruits have always played a very important role in this. Fruits can be consumed fresh in their natural form, even in winter (e.g. apples and oranges) when they are stored in the proper way. Before the 20th century, drinking squeezed fruit juices was the privilege of rich people. Nowadays, in order for us to consume some kinds of fruits all year round, it is necessary to produce concentrate with a high level of dissolved solids for storing the concentrate frozen, in the refrigerator or at room temperature, depending on the nature of the dissolved solid content. Concentrate can be diluted by water, or used to create a fruit juice with the same characteristics as the original fresh one, if the process of concentration has been carried out correctly. The industrial concentration of fruit juices is usually performed by
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382 Separation, extraction and concentration processes multistage vacuum evaporators. In most cases, the volatile components are recovered and added back into the concentrated product at a later stage. Evaporation often causes heat degradation of many valuable compounds. To prevent this significant decrease in quality, it is necessary to employ nonthermal methods of concentration, such as freeze concentration systems and membrane processes. The most common membrane techniques are the membrane filtration processes. Micro- (MF) and ultrafiltration (UF) are used as clarifying processes, whereas nanofiltration (NF) and reverse osmosis (RO) are used for pre-concentration of the juices (Lagana et al. 2000, Nene et al. 2002, Rektor et al. 2006, 2007).
13.2 Characteristics of foods/fluids in the fruit juice product sector In the world market there are countless fruit juice based products. They differ mainly in terms of raw material, composition, fruit content, nutritional value, sensory characteristics and packaging, but in some cases the biggest difference is the brand name (Hui et al., 2006). Fruit juice based drinks are classified on the basis of fruit content, into three categories: ∑ juices and fruit musts; ∑ fruit nectars; and ∑ soft drinks with fruit content. Juices and fruit musts are produced by mechanical procedures, mainly pressing, and the produced juice has the same taste, color and aroma as the original fruit. The final fruit juice composition is also identical to the original fruit. Juices cannot contain food additives (preservatives, aromas and coloring agents). Juices are consumed in their fresh form soon after production, or used as a raw material for concentration. Fruit juices can be divided into two subcategories: they can be filtered, i.e. clarified to be transparent (apple, grape), or they are cloudy, containing colloids and fibers like all citrus-based juices. Fruit nectars are made from fruit pulps or fruit juices diluted with sugar syrup. They usually contain only one fruit, such as apple, orange or peach, but they can be made from blends of more than one fruit juice or pulps. The preparation of blends and minimum fruit content are regulated by government standards, industrial specifications and other requirements (Hui et al., 2006). In order to ensure international trade, these standards conform to the Codex Alimentaria of the FAO/WHO Food Standard Program. The raw materials used in the production of fruit juices can be grouped as follows: citrus, pomaceous fruits, stone fruits, grape and berries. Some of the raw materials, such as, apple and orange are suitable for juice production
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Separation technologies in the processing of fruit juices 383 the produced juice can be consumed without any additive, but the juice of some berries (sea buckthorn, black and/or red currant) are very acidic and are enjoyable only when blended with sugar syrup or concentrate of apple juice or grape must.
13.3 Designing separation processes to optimize product quality in the fruit juice product sector Juice extraction, the removal of juice from fibrous solid particles, is the basic operation of fruit juice production. The fruit has to be prepared for this operation, i.e. the separation step. This preparation step depends on the type of fruit. In some instances, the fruit has to be chopped (apple, peach), whereas in other instances (cherry, sour cherry, plum, apricot) it is very important to remove the stem before chopping. A typical flow sheet of the natural juice and juice concentrate production is shown in Fig. 13.1. In this scheme of concentrate production, several conventional separation processes can be recognized: ∑ mechanical pressing of the juice from the fruit pulp; ∑ juice extraction from the marc by water as solvent; ∑ clarification of the fruit juice by centrifugation or filtration; ∑ concentration of the fruit juice by multistep vacuum evaporation. The fruit juice production technology presented in Fig. 13.1 is a typical one (Barta and Körmendy, 2007) for the production of fresh natural fruit juice and concentrate for longer storage. In the past few decades, some of the conventional separation processes have been replaced by newer ones, for example the clarification step, where the traditional method of clarification by centrifugal separation or filtration has been replaced by ultrafiltration, especially in the case of clarified or ‘transparent’ juice production, such as apple and grape (Barta and Körmendy, 2007). In this traditional technology, the juice yield can be improved by using enzyme treatment for pectin degradation, as well as through the extraction of the residual fruit juice from the marc using hot water (50–90 °C) (Hui et al., 2006). 13.3.1 Juice extraction by pressing For juice extraction by pressing, the liquid content of the fruits is separated from the solid particles. The most common method of this separation is a mechanical pressing of the juice out from the fruit pulp. The type of equipment utilized in this separation depends on fruit species. Where the hard parts of the fruit (e.g. the stem) have been previously removed (cherry, plum, apricot) it is possible to use typical mechanical pressing with higher pressure, whereas for pressing berry-type fruits, such as grapes and currants, mild pneumatic
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384 Separation, extraction and concentration processes Reception of the fruits
Cleaning of the washing water
Washing
Stem elimination Marc Juice extraction–pressing Secondary juice extraction
Aroma extraction
Clarification (centrifugation)
Enzyme treatment
Pasteurization
Clarification (centrifugation)
Storing
Concentration (evaporation)
Natural, unfiltered juice
Storing
Fruit juice concentrate
Fig. 13.1 Production of natural fruit juice and fruit juice concentrate using conventional separation processes.
pressing is more effective (Barta and Körmendy, 2007). In both instances, the pressing processs requires outside forces to create tension in the system and drain out the liquid, which results in some shape modification. In batch systems, the solid phase will stay in the pressing vessel, whereas the liquid will be drained out across a sieve and/or filter media. The remaining solid, with low liquid content, is called marc. The most important parameter of the pressing process is the liquid or ‘juice yield’, which refers to the percentage of juice pressed out, compared with the amount of raw material that was entered into the system. The juice yield is determined basically by the preparation and pretreatment (enzymatic hydrolysis or not) of the fruit before pressing, and the pressure applied (Barta and Körmendy, 2007; Hui et al., 2006). In the basket type batch pressing machines commonly used in the fruit juice processing industry, press machine volume is often decreased by mechanical or hydraulic forces; the space available for the fruit pulp/marc may also be
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Separation technologies in the processing of fruit juices 385 decreased by pneumatic interaction at lower pressures, up to a few bars, to press the juice out of the fruit. 13.3.2 Juice extraction using water as solvent Fruit juices can also be produced by solid–liquid extraction. This process can be characterized by the degree of extraction, expressing the amount of valuable substances extracted, compared with the total valuable matter content of the fruit. In this operation, mass transfer is controlled by the molecular diffusion of certain components, which can be improved by increasing the temperature of the operation. In order to increase the diffusion coefficient and the permeability of the cell walls, this solid–liquid extraction is performed at 50–70 °C. The mass transfer can be improved by increasing the specific area available for transport, in most instances by decreasing the particle size of the solid phase or by chopping the fruits. Another means of improving the mass transfer is increasing the concentration gradient for the mass transfer using a higher solvent–solid ratio, multistep cocurrent or countercurrent operation mode (Fonyó and Fábry, 1998). Diffusion juice extraction is usually carried out in double-screw extractor devices (Hui et al., 2006). 13.3.3 Juice clarification The extracted fruit juices are usually turbid, containing plant particles (fibers, cellulose, starch, and lipids) and colloids such as pectin, proteins, and polyphenols. Depending on the nature of the final product, these substances must be partially or totally removed to avoid further turbidity and precipitation and to improve sensory attributes such as taste, color, and odor. This clarification step can be performed by physicochemical or mechanical methods, as well as by using combinations of these methods. The physicochemical methods of clarification are those in which clarifying agents and enzymes are added during the procedure. For the clarification of fruit juices, mineral clarifying agents (bentonite and silicic acid), natural organic (gelatine) or organic polymers (polyvinypolypyrrolidone) are often used. The similarity between these clarifying agents is that all of them are charged (Barta and Körmendy, 2007), a characteristic which causes the bentonite to adsorb proteins, and the gelatine and polyvinylpolypyrrolidone to precipitate negatively charged particles (polyphenols and decomposed pectin). During juice clarification, the pectin molecules have to be decomposed, because they hinder aggregate formation and the settling of floating substances. Besides pectin hydrolysis, the decomposition of starch and proteins can be carried out on the same time scale using pectin, which results in the application of a mixture of enzymes (Hui et al., 2006). The aim of mechanical clarification is also to remove suspended solids using only mechanical separation processes like centrifugation or filtration. This separation process is performed in settling centrifuges, eliminating fibers
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386 Separation, extraction and concentration processes from cloudy juices. Filtration is the next step of fruit juice production. The traditional method is using devices based on a slurry layer and an operation mode using silica or perlite as additives. The filtration equipment should be frame filter press, continuous vacuum drum filter or others. Ultrafiltration and microfiltration have been widely applied in the clarification of filtered juice membrane filtration in the past two decades (Fonyó and Fábry, 1998; Bélafi Bakó, 2002). These membrane filtrations are able to solve the problem of clarification and filtration in one step. These clarified and filtered or cloudy juices are then ready for consumption. The juice can be packed and pasteurized and released to market, or it can be used as a suitable raw material for the production of semifinished products.
13.4 Production of fruit juice concentrate As shown in Fig. 13.1, after production (either by pressing or solid liquid extraction) and certain post treatments, the fruit juice can be concentrated for longer shelf life, decreasing the water content and increasing the total soluble solid (TSS) content. This concentration improves storage and transportation properties and costs, but it has to be carried out very carefully in order to avoid the loss of aroma components and valuable ingredients, and to minimize changes in the sensory properties of the juice. The most common method of juice concentration is evaporation (Barta and Körmendy, 2007). 13.4.1 Concentration by evaporation The most common method of fruit juice concentration is evaporation, and, from a physical point of view, this means the water evaporating from the boiling liquid phase. This process is carried out in devices such as an evaporator and steam provides the energy for boiling and water evaporation. However, valuable fruit components are mostly heat sensitive, so, for fruit juice concentrate production, vacuum evaporation is recommended (Ashurst, 2005). To decrease energy consumption during this process, batteries of 3–4 evaporator elements are commonly used (Fonyó and Fábry, 1998). The chemical, rheological and thermal characteristics of juices play an important role in this evaporation–condensation process. Because these parameters tend to vary from juice to juice, the operation parameters are different for different fruits or juices. The type of evaporators chosen for an individual juice is based on the characteristics of that particular juice. The most widely used type of evaporators are the film, pipe, plate, and centrifugal based ones (Barta and Körmendy, 2007). Evaporators are usually combined with aroma recovery units, in most cases a distillation column. The condensed aroma components are often remixed into the concentrate to improve the flavor,
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Separation technologies in the processing of fruit juices 387 or can be concentrated and applied as natural aroma extracts in other fruit products. 13.4.2 Concentration by freezing During freezing, the fruit juice is cooled down under water at 0 °C at atmospheric pressure. The water forms ice crystals and these ice crystals are separated from the suspension, thus lowering the water content, i.e. the concentration of TSS will be higher. Owing to the high cost of producing ‘cold energy’, this technology is only appropriate for the concentration of valuable and heat-sensitive fruit juices. This concentration method can be achieved in one or several steps, and, with this technology, the eutectic concentration of the fruit juice–water system can be achieved (Barta and Körmendy, 2007). The separation of the solid phase (ice) and the liquid phase (concentrated juice) should be carried out by mechanical separation: settling, filtration, and centrifugation. Only clarified juices should be concentrated using this technology, because for unclarified juice separation, the fibers and colloids are removed together with the ice. On the other hand, this type of concentration is a very gentle process, because there are no aroma, color or vitamin losses during this operation, owing to the low temperatures involved. Its disadvantages are high energy consumption and lower concentration efficiency compared with evaporative concentration (Barta and Körmendy, 2007). 13.4.3 Concentration by membrane separation processes The sterilization of fruit juice can be achieved without heating by using membrane separation processes, by the mechanical removal of microbes during clarification. On the other hand, water removal (concentration of fruit juice) can be achieved at or near room temperature (Jiao et al., 2004). For the clarification of the juice, ultrafiltration (UF) or microfiltration (MF) is suitable (De Carvalho et al., 2008; Cassano et al., 2006, 2007; He et al., 2007). Even when we use ultrafiltration we are removing more valuable components (polyphenols) than is ideal. For the concentration of fruit juices, reverse osmosis (RO) (Rektor et al., 2004; Kozak, 2005) and/or nanofiltration (NF) can be used (Kiss et al., 2004; Porter, 1990). The only problem is that during this concentration process on the retentate side of the membrane, the concentration of the TSS becomes higher and higher, causing higher osmotic pressure and decreasing the driving force of the concentration process, which is the difference between the transmembrane pressure (TMP) and osmotic pressure of the retentate (fruit juice concentrate) and permeate (practically water). Using the usual TMP (40–50 bar), the fruit juice can be concentrated up to 23–26% TSS (Belafi-Bako et al., 2000; Belafi-Bako, 2002). The advantages of this concentration technology include better quality juice concentrate as valuable components like aroma, vitamins, polyphenols are concentrated and
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388 Separation, extraction and concentration processes not destroyed, and, on the other hand lower energy consuption and lower cost in comparison with evaporation, as well as with freeze concentration (Ashurst, 2005; Kiss et al., 2004). Having used membrane technology to reach the final concentration (45–65% TSS) for storing the concentrate, at room temperature or little bit lower, the semi-product produced by reverse osmosis (23–26% TSS) should be further concentrated by membrane distillation (MD) (Laeson, 1997; Lagana et al., 2000; El-Bourawi et al., 2006) osmotic distillation (OD) (Lefebre, 1988; Courel et al., 2000; Thanedgunbaworn et al., 2007) or nanofiltration (NF) (Kiss et al., 2004; Porter, 1990; Vatai, 2007). From the previous statements, the fruit juice concentration by membrane technology could be solved in a three-step process: clarification (UF or MF), the production of ‘half-concentrate’ by a pressure-driven membrane process (RO), and the production of ‘final concentrate’ using three different methods (Vatai, 2007). One of the most valuable fruits in Hungary is the grape. As an example, the production technology for the must concentration is shown in Fig. 13.2, supported by experimental and optimization data (Vatai, 2007). The conditions for grape growing in the country are very good, and a large volume and good quality must is produced. Most of the vintage is processed by the viticulture and soft drinks industry. Because the ripening of the grape and the vintage happen just once a year, there can sometimes be a lack of grapes (or poor quality) or an overproduction. The preservation of grape juice can solve these problems. Viticulture can produce must concentrate from the surplus that can then be applied to the process of upgrading poorquality, low-sugar content grape juice, or the soft drinks industry can develop a new product for its customers. This technology for the production of must concentrate has advantages in terms of saving both energy and quality. However, it also has some other benefits in terms of food safety and the impact on the environment, because this is practically a clean technology:
Colloids, microorganisms
Must 12–16% TSS
MF
~Water
Clarified must 12–16% TSS
Concentrate 23–26% TSS
RO
NF
MD
OD
Concentrate 45–60% TSS
Concentrate 55–65% TSS
Concentrate 55–65% TSS
Fig. 13.2 Alternative methods of grape juice (must) concentration by complex membrane processes.
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Separation technologies in the processing of fruit juices 389 the main product is the must concentrate, the byproducts are the membrane filtered water and the concentrate of the clarification, which can be added to grape marc for ‘pálinka’ or spirit production, but can also be used as raw materials for the production of valuable products rich in antioxidant capacity by conventional or supercritical extraction. In the proposed must concentration technology, the first step is clarification by microfiltration. The aim of this step is: ∑
the removal of suspended solids, which causes a decrease in the viscosity and results in a higher filtration capacity, ∑ in further membrane filtration processes for must concentration, the clogging of the circulation channels, especially in spiral wound membrane filtration elements, is minimized, ∑ the micro-organisms are also removed, resulting in the sterilization of the must. Figure 13.3 contains typical experimental data of changing permeate flux during the clarification of typical Hungarian musts Kékfrankos and Furmint, by microfiltration. From Fig. 13.3 it can be seen that, after a certain period of time, the flux decreasing rate has been stabilized, reaching a steady-state flux. From the diagram it is also obvious that there are some differences between the permeate fluxes of white grape must (Furmint) and blue grape must (Kékfrankos), owing to the greater total solid content of the Kékfrankos must (Vatai, 2007). From the clarified must, the so-called ‘half concentrate’ can be produced by reverse osmosis, with TSS between 23 and 26%. This can be stored frozen for a long time, thus maintaining the valuable component concentration. The results of the concentration of Furmint must in a pilot scale RO plant in the laboratory of Food Engineering, Corvinus University of Budapest, 70 Furmint
Flux, J (L m–2 h–1)
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1.6
1.8
Fig. 13.3 Permeate flux changing during clarification of typical Hungarian musts, Furmint and Kékfrankos, by microfiltration.
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390 Separation, extraction and concentration processes is presented in Fig. 13.4. The flux of the permeate decreases during this process because of the lower driving force caused by the decreasing osmotic pressure of the concentrate, as well as by the increasing viscosity. From the diagram, it is also obvious that the final concentration at room temperature (50 bar transmembrane pressure and 600 L h–1 volumetric recirculation flow rate) was slightly more than 23% TSS and, at the end of the experiment, the permeate flux decreased almost to zero (Vatai, 2007). The ‘half-concentrate’ produced can be stored frozen for even longer but the storing costs, as well as the transportation costs of the semi-product from the producer to the drink producer, would be quite high if the concentration was produced near to the location where the fruit was grown. To optimize these costs, the next step of the must concentration process can be applied using conventional concentration technology (evaporation) or membrane technology, which has the benefits of a lower energy cost and better quality of concentrate (Jiao et al., 2004). The alternatives for further concentration procedures, shown in Fig. 13.2, are NF, because for NF the osmotic pressure on the permeate side is greater owing to a higher TSS concentration, causing a lower osmotic pressure difference, which results in a better driving force (Kiss et al., 2004; Porter, 1990). The loss of TSS in the permeate of NF should be compensated by returning the permeate of the nanofiltration to the feed of the reverse osmosis, as proposed by Porter (1990) for orange juice concentration. 24
10 9
23
8
Flux, J (kg m–2 h–1)
6
21
5 20
4 3
TSS (%)
22
7
19
2 18
1 0 0
1000
2000
3000
4000
5000 6000 Time (s)
7000
8000
17 9000 10000
Fig. 13.4 Concentration of Furmint must on a pilot-scale reverse osmosis plant in the laboratory of Food Engineering, Corvinus University of Budapest.
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Separation technologies in the processing of fruit juices 391 The second alternative for the production of the final concentrate is membrane distillation. The basic principle of this membrane separation process is the removal of water from the aqueous phase across a hydrophobic membrane when the membrane and the membrane pores are not wetted by the water phase on the feed and permeate side. The driving force of this separation process is the temperature, i.e. the vapor pressure difference on the two sides of the membrane. The separation process can be carried out in the hydrophobic hollow fiber membrane modules, characterized by a large specific mass transfer area of up to 10 000 m2 m–3. Membrane materials include polypropylene, polyvinylidenetetrafluoride, polytetrafluoroethylene (Lagana et al., 2000). The size of micropores is 0.2–1.0 mm. The porosity of the membrane is 60–80% of the volume, and the overall thickness is 80–250 mm, depending on the presence or absence of support. In general, the thinner the membrane and the greater the porosity of the membrane, the greater the flux rate. Experimental results of the concentration of typical Hungarian musts obtained by MD (Kékfrankos and Furmint), and the influence of the type of must, driving force and retentate concentration on permeate flux, are shown in Fig. 13.5. From Fig. 13.5, it is obvious that the influence of the type of must is not significant, but the driving force is the temperature difference, which has the main role in this separation process. Twice the temperature difference resulted in permeate fluxes three times higher, owing to the nonlinear behavior of the vapor pressure in relation to the temperature. OD is a recent membrane process (Lefebvre, 1988) (also known as osmotic evaporation, membrane evaporation, isothermal membrane distillation or gas membrane extraction) which has been successfully applied to the 2.0 Furmint DT 30 °C Kékfrankos DT 30 °C Furmint DT 15 °C Kékfrankos DT 15 °C
1.8
Flux, J (kg m–2 h–1)
1.6 1.4 1.2 1.0 0.8 0.6 0.4 0.2 0.0 10
20
30 40 50 Concentration of retentate CR (TSS%)
60
70
Fig. 13.5 Concentration of typical Hungarian musts Kékfrankos and Furmint, by membrane distillation, influence of the type of the must, driving force and retentate concentration on permeate flux.
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392 Separation, extraction and concentration processes concentration of liquid foods such as milk, fruit and vegetable juice, instant coffee and tea and various non-food aqueous solutions. This technique can be used to extract water from aqueous solutions selectively under atmospheric pressure and at room temperature, thus avoiding the thermal degradation of solutions. It is therefore particularly adapted to the concentration of heat-sensitive products such as fruit juices (Lefebvre, 1988). The process involves the use of a microporous hydrophobic membrane to separate two circulating aqueous solutions at different solute concentrations: a dilute solution and a hypertonic solution (usually salt solution). If the operating pressure is kept below the capillary penetration pressure of liquid into the pores, the membrane cannot be wetted by the solutions. The difference in solute concentrations and, consequently, in water activity of both solutions, generates, at the vapor–liquid interface, a vapor pressure difference causing vapor transfer from the dilute solution towards the stripping solution. The water transport through the membrane can be summarized in three steps: ∑ evaporation of water at the dilute vapor–liquid interface; ∑ diffusional or convective vapor transport through the membrane pore; ∑ condensation of water vapor at the membrane/brine interface. The typical OD process involves the use of a concentrated brine on the downstream side of the membrane as the stripping solution. A number of salts such as MgSO4, CaCl2, and K2HPO4 are suitable. Potassium salts of ortho- and pyrophosphoric acid offer several advantages, including low equivalent weight, high water solubility, steep positive temperature coefficients of solubility and safe use in foods and pharmaceuticals (Courel et al., 2000; Lefebvre, 1988). When compared with RO and MD process, the OD process has a potential advantage which might overcome the drawbacks of RO and MD for concentrating fruit juice, because RO suffers from the limitation of, high osmotic pressure whereas in MD some loss of volatile components and heat degradation may still occur owing to the heat requirement for the feed stream in order to maintain the water vapor pressure gradient. OD, on the other hand, does not suffer from any of the problems mentioned above when operated at room temperature. The most well-known module designed for OD is the Hoechst-Celanese Liqui-Cel membrane contactor with an effective area/ volume of 2930 m2 m–3, a maximum transmembrane differential pressure of 4.08 bar and a temperature operating range of 1–40 °C, and contains microporous polypropylene hollow fibers of Celgard membrane. These fibers are approximately 0.3 mm in external diameter, with a wall thickness of about 0.03 mm; they have a mean pore diameter of about 30 nm and a porosity of about 40%. The experiments of must concentration by osmotic distillation in the laboratories of Corvinus University of Budapest have been carried out using polypropylene hollow fiber modules with CaCl2 as osmotic solution. The
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Separation technologies in the processing of fruit juices 393 details of these experiments are described elesewhere (Kozak, 2005; Rektor et al., 2006). On a laboratory and pilot scale, the final must concentration was over 65 °Brix (Kozak et al., 2008; Rektor et al., 2006). One of the possibile means of reaching a final concentration of must which can be stored at room temperature is the integration of RO and NF. The first step of this concentration technology is concentration by RO but, as shown in Fig. 13.4, it can be used at least up to 23–26% TSS at usual transmembrane pressures (40–50 bar). Assuming that the transmembrane pressure should theoretically be higher in the case of RO and NF, that the experimental data collected at 40–50 bar transmembrane pressures can be used as the basis for the modeling, and that it can be extended to higher pressures, a theoretical design and optimization of a two-step must concentrator was carried out (Kiss et al., 2004; Vatai, 2007). The calculation of the optimal cut off between the RO and NF steps has been carried out by dynamic programming. The basis of these calculations was the total cost of the RO and NF step, as well as the cost of the complete concentration process. For this purpose, it is necessary to develop a relation between the cost and the concentration of the final concentrate. The simplest solution is to find a relation between the yield of the step (Y) and the cost of the concentration process as whole. Assuming, in this first calculation, that the permeate concentrations in comparison with the retentate concentration should be neglected, the equations can be written as:
YRO = (x1 – x0)/x1
where (cp ~ 0)
[13.1]
and in the case of NF:
YNF = (x2 – x1)/x2
where (cp ~ 0)
[13.2]
From the above equations, the relation between the total cost of a step, RO or NF can be written as:
∑ TCOST = ∑ ICOST + ∑ OCOST = f (Y)
[13.3]
where x0 is the feed must concentration in the RO step, x1 is the outlet must concentration from the RO step, and at the same time the inlet concentration of the NF step, and x2 is the final must concentration at the outlet of the NF, and at the same time the final concentration of the must at the outlet of the system of RO–NF concentrator. TCOST is the total cost of the must concentration technology, and for the RO and NF step, respectively, and ICOST is the investment cost of the must concentration technology, and for the RO and NF step, respectively, and OCOST is the operation cost of the must concentration technology, and for the RO and NF step, respectively. Figure 13.6 shows the changes in the total cost (Hungarian forints (HUF)/ year) of the two-step (RO–NF) must concentration plant, and the dependence upon the outlet concentration of the first (RO) [i.e. inlet concentration to
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394 Separation, extraction and concentration processes 1 560 000 1 540 000
Total cost (HUF/year)
1 520 000 1 500 000 1 480 000 1 460 000 1 440 000 1 420 000 1 400 000 1 380 000
15
20
25 30 35 Concentration of retentate (TSS%)
40
45
Fig. 13.6 Changes of the total cost (HUF/year) of the two-step (RO–NF) must concentration plant in relation to the outlet concentration of the first (RO), i.e. inlet concentration, to second (NF) step; using Hungarian prices from 2006 (1 EUR ª 250 HUF).
the second (NF) step], using Hungarian prices from 2006 (1 euro ª 250 HUF) (Vatai, 2007). From this diagram, it can be seen that the minimum system total cost is at 24% TSS concentration, which is the optimum for this two-step system. It means that the must should be concentrated from the feed to 24% TSS in the RO step, and that it then has to be turned to the NF step, when the final concentration of 45% TSS can be reached. These concentration processes should be carried out at theoretical transmembrane pressures of 100 and 110 bar, respectively.
13.5 References Ashurst P. R. (2005): Chemistry and technology of soft drinks and fruit juices, Blackwell Publishing Ltd., UK. Barta J., Körmendy I. (2007): Basic processes of vegetable and fruit processing technologies (In Hungarian), Mezőgazda Kiadó, Budapest. Bélafi-Bakó K., Gubicza L., Mulder M. (2000): Integration of membrane processes into bioconversions, Kluwer Academic/Plenum Publishers, New York, USA. Bélafi-Bakó K. (2002): Membrane separation processes (in Hungarian), Veszprémi Egyetemi Kiadó, Veszprém. Cassano A., Tasselli F., Drioli E. (2007): Ultrafiltration of kiwifruit juice using modified poly(ether ether ketone) hollow fibre membranes, Separation and Purification Technology 57, 94–102. Cassano A., Marchio M., Drioli E. (2006): Clarification of blood orange juice by ultrafiltration: analyses of operating parameters, membrane fouling and juice quality, Desalination 212, 15–27. © Woodhead Publishing Limited, 2010
Separation technologies in the processing of fruit juices 395 Courel M., Dornier M., Herry J. M., Rios G. M., Reynes M. (2000): Effect of operating conditions on water transport during the concentration of sucrose solutions by osmotic distillation. Journal of Membrane Science 170, 281–289. De Carvalho L. M. J., de Castro I. M., Bento da Silva C. A. (2008): A study of retention of sugars in the process of clarification of pineapple juice (Ananas comosus, L. Merril) by micro- and ultra-filtration, Journal of Food Engineering 87, 447–454. El-Bourawi M. S., Ding Z., Ma R., Khayet M. (2006): A framework for better understanding membrane distillation separation process, Journal of Membrane Science 285, 4–29. Fonyó Z., Fábry G. (1998): Basic principles of unit operations (in Hungarian). Nemzeti Tankönyvkiadó, Budapest. He Y., Ji Z., Li S. (2007): Effective clarification of apple juice using membrane filtration without enzyme and pasteurization pretreatment, Separation and Purification Technology 57, 366–373. Hui Y. H., Barta J., Cano M. P., Gusek T. W., Sidhu J. W., Sinha N. Editors (2006): Handbook of fruits and fruit processing, Blackwell Publishing. Jiao B., Cassano A., Drioli E. (2004): Recent advances on membrane processes for the concentration of fruit juices. Journal of Food Engineering 63, 303–324. Kiss I., Vatai G., Bekassy-Molnar E. (2004): Must concentration using membrane technology. Desalination 162, 295–300. Kozák Á., Bánvölgyi S., Vincze I., Kiss I., Békássy-Molnár E., Vatai G. (2008): Comparison of integrated large scale and laboratory scale membrane processes for the production of black currant juice concentrate, Chemical Engineering and Processing 47, 1171–1177. Kozák Á. (2005): Experimental investigation of must concentration by reverse osmosis and membrane distillation (in Hungarian). MSc thesis, Corvinus University of Budapest. Laeson K. W., Lloyd D. R. (1997): Membrane distillation, Journal of Membrane Science 124, 1–25. Lagana F., Barbieri G., Drioli E. (2000): Direct contact membrane distillation: modelling and concentration experiments, Journal of Membrane Science 166, 1–11. Lefebvre M. S. M. (1988): Method of performing osmotic distillation. US Patent 4,781,837, 1 November. Nene S., Kaur S., Sumod K., Joshi B., Raghavaro K. S. M. S. (2002): Membrane distillation for the concentration of raw cane-sugar syrup and membrane clarified juice, Desalination 147, 157–160. Porter M. C. (l990): Handbook of industrial membrane technology, Noyes Data, Park Ridge Rautenbach R. (1997): Membranverfahren, Verlag, Berlin Rektor A., Kozak A., Vatai G., Bekassy-Molnar E. (2007): Pilot plant RO-filtration of grape juice, Separation and Purification Technology 57, 473–475. Rektor A., Pap N., Kókai Z., Szabó R., Vatai G., Bekassy-Molnar E. (2004): Application of membrane filtration methods for must processing and preservation. Desalination 162, 271–277. Rektor A., Vatai G., Békássy-Molnár E. (2006): Multi-step membrane processes for the concentration of grape juice, Desalination 191, 446–453. Thanedgunbaworn R., Jiraratananon R., Nguyen M. H. (2007): Mass and heat transfer analysis in fructose concentration by osmotic distillation process using hollow fibre module, Journal of Food Engineering, 78, 126–135. Vatai G. (2007): Processing of agricultural raw materials and foods by complex membrane separation processes (in Hungarian), Doctoral thesis, Hungarian Academy of Sciences, Budapest.
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396 Separation, extraction and concentration processes
14 Separation technologies in oilseed processing M. A. Williams, Anderson International Corp., USA
Abstract: The two procedures used to separate oil from oilseeds: crushing in mechanical screw presses or extraction with solvent, are described. Optimization of the various preparation steps required for both procedures is discussed. Processing steps for satisfactory recovery of solvent in solvent extraction plants are explored. Some of the currently available equipment for mechanical crushing and for solvent extraction is described. Key words: separation, solvent extraction, oilseeds, screw press.
14.1 Introduction Vegetable oils have been separated from oilseeds for many centuries, starting with relatively simple procedures for easily processed oilseeds such as sesame, peanut, and oil from oil-saturated fruits, such as olives. Gradually, over the years, more sophisticated procedures and better equipment led to the crushing of many oilseeds that did not easily yield their oils. These procedures grew into the unit operations of cleaning to remove dirt and trash, dehulling to remove low-oil/high-fiber components, flaking to provide uniform-thickness particles, and cooking to harden some seed proteins that are too soft to withstand the high pressure exerted by modern screw presses. Some oilseeds, such as cottonseed, require very elaborate preparation for most efficient oil yields. Equipment serving all the unit operations, in particular screw presses, became more durable and more efficient. In addition to mechanical screw pressing, solvent extraction became a method to obtain vegetable oils, and the solvent extraction systems involved their own unit operations and specialized equipment. Today, modern high© Woodhead Publishing Limited, 2010
Separation technologies in oilseed processing 397 pressure mechanical screw press systems and solvent extraction systems operate at high capacity and high efficiency. Small-volume processors prefer mechanical screw pressing to avoid erecting costly solvent extraction systems. High-volume processors prefer solvent extraction to avoid having multiple mechanical screw presses operating in parallel. In this chapter, procedures and equipment used in obtaining oils from oilseeds will be described, mentioning early systems, elaborating on modern systems, and focusing on newer innovations in mechanical screw pressing and solvent extraction.
14.2 Preparation for oilseed processing Typical preparation involves cleaning, dehulling, flaking, and cooking. Some oilseeds require all steps, some only a few steps. 14.2.1 Cleaning Incoming oilseeds are cleaned to remove sticks, pieces of stems and leaves, sand and dirt. Vibrating screens suffice for most oilseeds, but sometimes pneumatic aspiration chambers are also used to separate light-weight impurities from the heavier seeds. Some oilseeds contain appreciable amounts of weed seed, which can be screened out and sold as bird seed. Magnets placed before easily damaged equipment remove tramp iron to protect sensitive equipment from damage. 14.2.2 Dehulling Hulls should normally be removed from oilseeds before screw pressing. Hulls contain low levels of oil. Any hulls entering the screw press with the kernels (or meats) absorb some of the oil released from the kernels, thereby reducing oil yield. Hulls also reduce screw press capacity because a greater volume of material is passing through the screw press. Hulls are high in fiber and are more abrasive, causing greater wear on the processing equipment and greater consumption of horsepower. Dehulling, particularly of soybean, can be done ‘hot’, ‘warm’, or ‘cold’. The soybean seeds are cracked using hammer mills with swinging or fixed hammers or in corrugated breaking rolls, and the cracked seeds are passed through a chamber against a counter-current flow of air, which removes the hulls that have been loosened in the cracking step. Other seeds are dehulled in machines of two basic designs: bar hullers and disk hullers. Larger seeds surrounded by thick, hard shells are dehulled and separated by various methods, sometimes even by hand. Cottonseed, because it comes into the processing plant covered with short cotton fibers (lint), is first delinted using saw delinters, sometimes going through more than one set of delinters. Fibers removed by the different ‘cuts’ find a ready market as cotton batting or in cellulose manufacturing. © Woodhead Publishing Limited, 2010
398 Separation, extraction and concentration processes 14.2.3 Flaking Oilseed kernels or meats are usually cracked into smaller particles and then flattened into thin flakes, if the meats are to be cooked before screw pressing or sent directly to a solvent extractor. This helps to ensure a more uniform cook before screw pressing. Flakes of uniform thickness also provide for a more uniform and more efficient penetration of solvent in solvent extractors. Flaking is usually done with the meats at approximately 10% moisture and a temperature of around 70 °C (158 °F). At that moisture and temperature, the oilseed meats easily smear into flat flakes without shattering into fine particles. Two types of flakers are used. One type employs two heavy steel rolls in parallel, almost touching each other, and revolving in opposite direction, so that both rolls pull the meats down between them. Mechanical jack screws or hydraulic cylinders apply pressure against the bearing blocks supporting the rolls so that the rolls do not spread apart when the meats pass between them. The second type consists of a set of flaking rolls stacked on top of each other, usually five high. This has an advantage in that the weight of the upper rolls provides pressure on the lower rolls to make thin flakes. The meats flow down the stack of rolls and are progressively flattened into thinner and thinner flakes. 14.2.4 Cooking Many oilseeds have soft meats, which are even softer when the hulls are removed. Such dehulled meats can easily form an oily paste within the screw press. These meats are usually cooked to stiffen and harden the protein so it can better withstand the pressures generated by the screw press. Cooking is also used to deal with undesired heat-sensitive components such as undesirable enzymes in some oilseeds and with gossypol in cottonseed. Gossypol can impart a red color to the oil and render the deoiled solids unfit for some animal feeds. Cooking can combine the gossypol with the protein in the solids preventing the gossypol from being extracted with the oil. If the combination is strong enough, it will reduce the gossypol’s toxic effect in defatted solids that go into animal feeds. Two types of cookers are used to cook oilseeds: ‘stack cookers’ and ‘horizontal cookers’. A stack cooker is comprised of steam-heated chambers stacked in a vertical housing. Steam jackets on the bottom surface of each stack supplies the heat for cooking. A horizontal cooker consists of a bank of two or more cylindrical vessels positioned horizontally, usually one above the other. Both horizontal vessels have the outer walls steam jacketed. An advantage of horizontal cookers is that raw feed drops into a large volume of tumbling material that is already hot; thereby quickly coming up to desired cooking temperature. In a stack cooker, on the other hand, the raw feed falls on top of a thick layer of slowly agitated material that is agitated
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Separation technologies in oilseed processing 399 beneath the surface, causing some delay in raw material reaching desired cooking temperature. If the raw material contains an enzyme that damages oil quality, the enzyme can become active long enough inside the top stack to do its damage before the temperature becomes high enough to inactivate the enzyme. This long delay in reaching the desired cooking temperature does not happen in a horizontal cooker. There is a slight possibility, however, that some undercooked oilseed could move across the surface of the oilseed in the cooker instead of being churned down into lower levels within the cooker. If this happens, then some undercooked oilseed could move quickly toward the discharge port, leaving the cooker before it is fully cooked. If an enzyme is present, and if it is critical that the entire oilseed is exposed to adequate retention time, it would be prudent to use two horizontal cooking vessels in series before sending the oilseed on toward downstream equipment. Cooking is done at elevated moisture, usually 9 to 12%, which is too high a moisture level for efficient pressing. Cookers are followed by similar vessels that serve as dryers to bring the moisture down to 2–5%. This is done in vented vessels, usually with a draft of air passing through them, and with a heat source, usually steam jackets, to elevate the oilseed temperature high enough to drive the moisture out, usually 104–120 °C (219–248 °F).
14.3 Extrusion preparation for oilseed processing 14.3.1 Origin of extrusion Extrusion cooking expanders were accepted into pet food manufacture in 1954. These machines cooked and puffed cereal grains and nutritionally balanced feed formulations for the petfood industry. In the 1960s, cooking extruders were applied to the preparation of rice bran ahead of solvent extraction. In the 1970s, extrusion was also used to form porous collets from flaked soybean and other oilseeds. Extrusion brings the incoming oilseed to desired moisture, temperature, and pressure conditions within 10 s. If there are troublesome enzymes present, extrusion can bring the oilseed enzymes to inactivation conditions long before the enzymes have time to do damage. Extrusion is even more effective to inactivate enzymes than the horizontal, atmospheric pressure cookers described earlier. Extrusion has been successfully employed for inactivating enzymes such as lipase in rice bran (Williams and Baer, 1965) and urease in soybean (Williams, 1991). 14.3.2 Extrusion before solvent extraction The first application for extrusion before solvent extraction was rice bran (Baer, 1966). Rice bran arrives at the plant as a fine powder, which presents two problems: the fine powder is very difficult to process because solvent © Woodhead Publishing Limited, 2010
400 Separation, extraction and concentration processes cannot percolate through a bed of fine powder. Even more troublesome, the bran contains the enzyme, lipase, which splits vegetable oils rapidly into free fatty acids. Activated as soon as the bran is exposed to air, lipase will raise the free fatty acid level by approximately 3 to 7% every day until the level reaches 50–75%. Cooking conditions during extrusion quickly inactivates lipase and converts the bran into a coarse meal that permits rapid percolation of solvent through the bed of material within the extractor. Similar extrusion also transforms flaked soybean and flaked cottonseed into porous collets that handle better in a solvent extractor. Farnsworth described extrusion of cottonseed (Farnsworth et al., 1986). 14.3.3 Closed-wall extruders An extruder consists of a rotating wormshaft within a cylindrical barrel (Fig. 14.1). Material to be extruded is fed into one end of the barrel, and the wormshaft forces it out through a die plate at the discharge end. The wormshaft flighting is not a continuous wrap; it is composed of partial wraps and unflighted segments. Stationary pins protruding from the barrel wall intermesh between the individual wraps. The combination of rotating flights and stationary pins masticates the oilseed, quickly blending injected steam into the oilseed to elevate the moisture and temperature within the extruder.
Fig. 14.1 Typical closed-wall extruder (courtesy Anderson International Corp.).
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Separation technologies in oilseed processing 401 The oilseed within the extruder is subjected to high pressure, 1379–4137 kPa (200–600 psi), and is pressure-cooked at optimum moisture and temperature to convert the protein into a tacky, elastic‑like condition. The injected steam releases its heat of vaporization, which helps to heat the oilseed. Additional heat is generated by friction from the rotating shaft. 14.3.4 Slotted-wall extruders Oilseeds containing more oil than cottonseed meats (30–33%) are difficult to process through a closed-wall extruder because the liberated oil accumulates within the extruder and disturbs the steady‑state operation. Also, the extruded collets cannot reabsorb all the liberated oil, and some oil is lost in product transport. If this is troublesome in extrusion plants, sometimes deoiled meal is blended with the incoming oilseed to dilute the oil level. This corrects the problem of excess liberated oil, but the recycling of meal increases the work done by the extruder, the solvent extractor, and the meal desolventizer, and reduces the total overall capacity of the oil mill. Anderson International (Anderson) introduced a slotted-wall extruder that permits a controlled release of excess oil through the slotted wall and produces collets at 20–30% oil (Williams, 1990) (Fig. 14.2). This slotted-wall extruder can process full fat safflower (at 42% oil), sunflower (at 42–44% oil), and peanut (at 45% oil) producing collets at 20–30% oil for solvent extraction. The liberated oil comes through the slotted wall and thereby
Fig. 14.2 Typical slotted-wall extruder (courtesy Anderson International Corp.).
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402 Separation, extraction and concentration processes bypasses the solvent extractor. Typical preparation before extrusion is to crack to approximately 1.6 mm (1/16≤) particles or flake to 0.3 mm (0.012≤); heat to approximately 60–71.1 °C (140–160 °F); and reduce moisture to approximately 8% without cooking the protein. Closed-wall and slotted-wall extruders are also fitted with discharge cone chokes rather than die plates at the discharge end. The cone is mushroom shaped and is moved laterally toward and away from a matching socket mounted on the extruder’s discharge. A hydraulic cylinder moves the cone in and out. Pressure generated by the extruder causes the oilseed to bear against the cone. Hydraulic fluid trapped within the cylinder holds the cone in a fixed position. Extruded product flowing across the cone into atmospheric conditions ‘expands’ with internal pores similar to the way collets expand when exiting the dies in a die plate. The sheets of porous cake falling past the cone break up in transit similarly to the way collets break up. 14.3.5 Dry extrusion Extrusion is also used to prepare oilseeds for subsequent screw pressing. Nelson et al. (1987) processed coarsely ground soybean through a high shear, closed-wall extruder, thus obtaining a hot, foamy mixture of small particles, liberated oil, and vaporizing moisture. The product upon discharge acted and looked like a fluid owing to moisture boiling through the freshly liberated oil. The soybean was introduced into the extruder at ambient temperature and moisture. The extruded product exited at 135 °C (275 °F), and the excess heat caused the moisture to flash down to 6–7% moisture. The heat is generated by friction through a specific configuration of screws, sometimes including steam locks, and by a cone attached to the end of the rotating shaft positioned into a stationary conical socket attached to the discharge plate. The contour of the socket matches the rotating cone and has a circular orifice in the center of the socket through which the extrudate exits. Shear is influenced by the proximity of the rotating cone to the stationary socket. 14.3.6 Dox-Hivex™ extruders Extruders adapted for dry extrusion, both closed‑wall and slotted‑wall, can be fitted with adjustable jaws or a manually positioned plunger to apply a choking action against the discharging oilseed (Fig. 14.3). These extruders are usually operated at low moisture so as to provide for high-shear rupturing of the oil cells. With little or no preparation, oilseeds are converted to a foamy, semi‑fluid extrudate at 121–148.9 °C (250–300 °F) that, as it cools, flashes down to 5–7% moisture. Shear is also influenced by the proximity of the cone point to the discharge port and the jaws and by positioning the jaws relative to each other or, if a plunger is used, by positioning the plunger relative to the discharge port. The plunger, a recent improvement replacing
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Separation technologies in oilseed processing 403
Fig. 14.3 Dox-Hivex™ extruder (courtesy Anderson International Corp.).
the jaws with a manually positioned plunger, can be moved close enough to the discharge port to almost completely block any discharge of material. 14.3.7 New developments in extrusion Extrusion preparation ahead of full-pressing has proven to be much more effective than traditional preparation. Two screw press manufacturing companies (that also manufacture extruders), Anderson International Corp and Insta-Pro, have developed extrusion preparation procedures pioneered by Nelson et al. (1987).
14.4 Mechanical pressing of oilseeds A mechanical screw press (Fig. 14.4) accepts a continuous stream of oilbearing material, compresses it under very high pressure exerted by a wormshaft rotating within a slotted-wall cage, the flights on the wormshaft propelling the oilseed forward as the shaft exerts pressure. The pressure releases the oil, most of which flows out through the slotted-wall cage. The deoiled solids flow through an adjustable port, like a movable cone, at the discharge end of the press (Fig. 14.5). The most common screw press application is full-pressing. Oilseeds are processed to liberate as much oil as possible through the full-press. Residual oil levels are usually measured as ether-extract using AOCS
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404 Separation, extraction and concentration processes
Fig. 14.4 Typical screw press (courtesy Anderson International Corp.).
Fig. 14.5 Cone choke on screw press (courtesy Anderson International Corp.).
method Ba 3-38 (or other equivalent and industry accepted procedures). A full-press generates an internal pressure of around 96 500 kPa (14 000 lbs in–2) so as to press out as much oil as possible. Earlier small-capacity full-presses reduced residual oil levels to between 3 and 5% in the exiting solids. Currently, higher capacity full-presses produce residual oil levels of 5–8%. With minor modifications, the same presses are (and always have been) used to partially press oilseeds as preparation in front of subsequent
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Separation technologies in oilseed processing 405 full-presses or as preparation for solvent extraction. The objective in prepressing is to arrive at 15–25% residual oil and run at higher capacities consuming less horsepower. Oil liberated by a screw press contains around 2–5% solids, but can contain up to 15% solids. The solids are removed in two steps. Step one employs a rectangular settling chamber equipped with a drag assembly that drags settled solids from the floor of the chamber, up an end wall, and over a wedge-wire drainage screen so that the solids can free drain as much as possible before recombining with fresh material going to the screw press. Step two clarifies the oil in a horizontal pressure leaf filter. Centrifuges, especially if they are three-phase (introducing water to flush solids out of the oil), are not well received in oilseed press plants because centrifuged solids, if reintroduced into the feed stream, are more difficult to press than solids obtained from filter presses. The filter press solids drop into a hopper equipped with a variable speed discharge screw that meters a stream of filter cake to blend in with fresh material going to the press. The feed to the screw press should be a homogeneous blend of fresh material, drainage chamber solids, and filter press cake. If the ratios of the three components are kept constant, the screw press will operate in a steady-state mode. 14.4.1 New developments in screw pressing There is a growing interest, today, in obtaining vegetable oils for food uses without the use of chemical solvents. The long-standing traditional method of achieving this is screw pressing; and screw pressing has been in use for more than a century. Since their introduction, screw press requirements have changed considerably from a time when labor was inexpensive and capital equipment costs were of concern. Today, labor costs are very high, and any simplification of time-consuming maintenance procedures would, of course, make mechanical screw presses much more attractive. A case in point is the tedious process of putting new barrel bars and spacing clips into the drainage cages. Historically, drainage cages are constructed of two cage halves that are clamped tightly together to surround the worm shaft propelling the oilseed through the screw press (Fig. 14.6). Barrels bars are assembled into the cage half, by placing a cage half horizontally on a work stand and manually positioning barrel bars and thin metal spacing clips of various thicknesses (Fig. 14.7). The spacing clips are used to provide gaps between the bars where oil can escape from the cage. Because very high pressure exists within the cage, the barrel bars must be captured tightly within the supporting structure of the cages to prevent any shifting of the bars during operation. This is a very time-consuming and labor-intensive task. However, if the assembly is not done correctly, the screw press will perform poorly. A desirable improvement is for screw press manufacturers to provide cages
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406 Separation, extraction and concentration processes
Fig. 14.6 Set of cages and shaft (courtesy Anderson International Corp.).
Fig. 14.7 Placing spacers (courtesy Anderson International Corp.).
that are easier to assemble and do not require so much precision in assembly. All screw press manufacturers are looking at this. The cylindrical drainage cage is constructed in two longitudinal halves adjacent to the wormshaft, sections are bolted tightly together to withstand © Woodhead Publishing Limited, 2010
Separation technologies in oilseed processing 407 the pressure. Inserted between the two cage sections are two lug bars called ‘knife bars’, 180° apart, running the entire length of the cage. Both knife bars have lugs protruding through the channel and intermeshing with the individual worm flights. The shaft is either of one piece construction or is made of individual worm segments slid onto a central shaft and keyed into position. Wormshafts are configured to progressively increase compaction on the oilseed as it is pushed from worm segment to worm segment. This ensures that compaction does not diminish as the volume of the oilseed is reduced owing to both the compression of the solids and the escape of noncompressible oil. This is done by reducing channel depth (the open space between the inside diameter of the barrel and the hub surface of the central shaft) and by decreasing the pitch of successive worm flights. Sometimes both techniques are employed. Some screw presses employ a force feeder to ensure sufficient inlet capacity to achieve desired screw press capacity. This introduces a second wormshaft and barrel plus a second drive mechanism to power the force feeder. The force feeders on large capacity screw presses are always driven by a smaller horsepower motor than the main wormshaft. Sometimes the smaller drive motor of the forced feeder reaches full horsepower consumption when the main drive motor is partially loaded. More input cannot be introduced because the force feeder will overload, but the main drive motor could easily accommodate greater capacity. Some screw press manufacturers are designing screw presses with a single motor driving both shafts or a gravity fed inlet not requiring a separate drive. A single drive motor allows for maximum available horsepower consumption on the main motor and does away with the problem of balancing the motor loads between the two drive motors. Some screw press manufacturers are working on new technology to improve capacity, reduce horsepower consumption and, by reducing horsepower consumption, reduce wear. Studies are also underway to characterize by mathematical modeling how a wormshaft configuration (profiles of pitch, channel depth, and thickness of flighting) influences the progressive compaction of the oilseed as it is conveyed by the wormshaft through the drainage cage. Most screw press manufacturers have shaft profiles that they know from experience have given satisfactory performance in the past. Currently, with the requirement of high capacity production plants, especially with concern of producing products that have not been exposed to chemicals such as hexane, there is interest in very high capacity screw presses that require less connected horsepower. Target capacities of 150 Mt d–1, or more, pressing to 3–5% residual oil using a smaller than traditional size motor is desirable, but no full-press offered today can achieve that performance. Screw press manufacturers are looking at mathematical modeling of shaft profiles to assist in more optimum shaft design. The objective is to design shaft profiles that press out maximum oil and dissipate minimum energy into friction so as to achieve maximum capacity at minimum horsepower consumption.
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408 Separation, extraction and concentration processes Most manufacturers are not going to give away their trade secrets, but some of the results of these efforts have already been published. Anderson International published some of their work in 1983 (Williams, 1983). Vadke et al. (1988) of the University of Saskatoon, developed a mathematical model of screw press operation by superimposing filtration analysis onto screw press theory. Predictions of how changes in shaft rotation and choke opening would affect screw press performance obtained by the mathematical model agreed reasonably well with experimental results obtained using a small laboratory screw press. Screw press oil yields can be increased by using supercritical CO2 to assist in obtaining the oil from a traditional screw press. Experiments in the early 1980s showed that significant increase in oil yields and in oil quality can be achieved by using supercritical CO2 (Friedrich et al., 1982) and (Friedrich and Pryde, 1984). Supercritical CO2 injected into the oilseed within the screw press cage can leech through the oilseed, dissolving the oil, and flow out with the oil through the drainage cage. Voges et al. (2008), reported oil yields increasing from 27% to as much as 71% when supercritical CO2 is used. A joint venture between two manufacturers is offering systems to do this (14.4.2). 14.4.2 Full-press suppliers There are many suppliers of small capacity screw presses offered for niche markets of special oils and small biodiesel plants. Vegetable oil production of oils for food purposes, however, is generally done in high capacity production plants where high capacity screw presses are desired. A list of well-known suppliers of high capacity screw presses and solvent extractors is available (Anon, 2009b). These suppliers also offer auxiliary equipment, and some of them offer entire plants. Anderson International has offered screw presses for many years. Their presses have been very well received in the past, but newer demands for screw presses favor higher capacity, less horsepower consumption, and simpler maintenance. Anderson has been working on a full-press of high capacity, consisting of a single motorized wormshaft and containing a simpler, less expensive choke (Fig. 14.5). One model, the Victor600™ (Fig. 14.8), is currently in field operation; another larger model is planned. The Victor600™ having a 6≤ (152 mm) diameter bore (inside barrel diameter) is rated at 50–75 Mt d–1 capacity, and is operating on sunflower and canola. The Victor1200™ with a 12≤ (304 mm) diameter bore is rated for 150 Mt d–1. Both presses feature a new choke design that permits replacing worn choke parts without disassembly of the wormshaft or removal of the drainage cages. Anderson also manufactures extruders and has developed an extrusion preparation system employing an extruder with a drainage cage followed by a full-press to press out the remaining oil. This ‘combo’ system (Fig. 14.9), utilizes a slotted-wall extruder for preparation, which can significantly increase the capacity of the full-press.
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Separation technologies in oilseed processing 409
Fig. 14.8 Victor 600™ press (Courtesy Anderson International Corp.).
Desmet Ballestra offers their line of Sterling Series presses (Fig. 14.10). These screw presses have 8≤ (203 mm) bores up to 13.75≤ (349 mm). Desmet purchased Rosedown several years ago, and some of the Sterling presses are still referred to as Rosedown Presses (www.rosedowns.co.uk). Sterling presses feature ‘multi-stage’, low-compression worm assemblies, and have an internal cake breaker. They can also be equipped with a ‘barring motor’, which is used to empty the press after power outages without requiring opening of the cages. Desmet is a major supplier of biodiesel plants around the world and has partnered with OriginOil (http://www.originoil.com/about-us/company.html) of Los Angeles, CA to commercialize OriginOil’s patent-pending process to extract oil from algae (http://www.originoil.com/company-news/originoilfiles-international-patent-for-low-energy-high-efficiency-algae-production. html). OriginOil’s patent is for a method to grow algae and provide sufficient light and CO2 for rapid growth, a system for rupturing oil cells, and for producing electricity with closed-loop CO2 recycling. The new process does not require drying the algae, which provides a substantial savings in electrical and thermal energy costs over traditional processing http://www. originoil.com/company-news/originoil-announces-partnership-agreementwith-desmet-ballestra-at-naa-conference.html. Dupps makes five models of Pressor® with capacities of 1500 to 12 000 lbs h–1 (18–144 Mt d–1). Bore diameters are 7≤ to 13≤ (177–330 mm) (Fig. 14.11). Pressors have hydraulically operated chokes that can automatically exert uniform discharge pressure by means of a controlled hydraulic pressure maintained on the hydraulic cylinder. The choke plunger ‘floats’ in equilibrium between the dynamic forces generated by the full-press and the force exerted by the hydraulic cylinder. When hydraulic pressure is held constant, the © Woodhead Publishing Limited, 2010
410 Separation, extraction and concentration processes Cottonseed Steam
Conditioner
Steam vapors
200 TPD Cottonseed meats
Condensate
Dox/Hivex™ extruder
Steam vapors
Expanded meal 175 TPD Vegetable oil to filtration 25 TPD
Steam
Dryer
Condensate
Expanded meal 167.4 TPD
11D expeller®
Total oil to filtration Hivex: 25 TPD Expeller: 26.6 TPD Total: 51.6 TPD
Vegetable oil to filtration 26.6 TPD
Press cake at 6% R.O. to meal grinding 140.8 TPD
Fig. 14.9 Extrusion–screw press system (courtesy Anderson International Corp.). TPD, metric tons per day; RO, residual oil.
plunger ‘floats’ in and out to ensure that the pressure within the screw press remains as steady as the oil pressure within the hydraulic cylinder, which pressure being controlled, does remains steady. © Woodhead Publishing Limited, 2010
Separation technologies in oilseed processing 411
Fig. 14.10 Desmet Ballestra press (courtesy Desmet Ballestra).
Fig. 14.11 Dupps screw press (courtesy The Dupps Company).
French Oil Machinery (Fig. 14.12) offers a line of Achiever presses; six models for full-pressing, which can also be used for pre-pressing before solvent extraction. Capacities range from 11 to 136 Mt d–1 on full-pressing. Achiever presses have force feeders, water-cooled drainage cages as well as water-cooled main worm shafts, and a cone choke to create back pressure against forward flow of material through the press. Harburg-Freudenberger offers several models of presses for full-pressing and pre-pressing. The largest full-press is Model EP-22 (Fig. 14.13); rated for 100–120 Mt d–1 and equipped with a 250–400 kW (339–543 hp) motor. Residual oils are 5–8%. Choking is accomplished through a fixed throttle
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412 Separation, extraction and concentration processes
Fig. 14.12 French OM press (courtesy The French Oil Machinery Company).
Fig. 14.13 Harburg-Freudenberger press (courtesy Harburg-Freudenberger).
ring that replaces mechanically adjustable chokes. Harburg-Freudenberger has also developed a worm design based on pressure worms to build up pressure alternating with pressure equalizing and relaxation via internal conical throttle rings. Harburg-Freudenberger has entered into a joint venture agreement with Crown Iron Works to offer Harburg-Freudenberger screw presses employing liquid CO2 injection to improve release of oil (the Hiplex™ process). Screw presses using this technology can reduce residual oils from 5–8% to 3–5%. Safe Soy Technologies, Elsworth, Iowa, recently put a Hiplex™ system on stream on soybean (Radio Iowa, 2008). Experimental work elsewhere in Hamburg University of Technology, reports significant increase in oil yields from soybean and rapeseed by injecting CO2 into a screw press (Voges et al., 2008). Willems P, Kuipers N and de Haan A, of University of Twente,
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Separation technologies in oilseed processing 413 working with sesame, linseed, rapeseed, palm kernel, and jatropha, reported oil yields 30% improved over conventional screw pressing (Willems, 2007), and claimed that displacement of oil by dissolved CO2 was the major cause of increased oil yields. The CO2 can be recovered, but recovery is expensive, and sometimes, in smaller capacity plants, the value of the extra oil does not justify the added cost to recover CO2. Insta-Pro International offers several models of extruders for preparation of oilseeds before pressing, the two largest extruders being Model 9400 with 300 connected horsepower (224 kW), and the Model 2000 Double Barrel with two 150 hp (112 kW) motors (Fig. 14.14), both models capable of 6720–7680 lbs h–1 (3056–3490 kg h–1) capacity. The extruded oilseed can then pass into several models of Insta-Pro screw presses, the largest being Model 5005, with a 60 hp (44 kW) motor, which can achieve 4000–4400 lbs h–1 (1818–2000 kg h–1) capacity (Fig. 14.15). 14.4.3 Other presses Duyvis Wiener BV has been manufacturing continuous pot presses for obtaining cocoa butter from cocoa bean (Fig. 14.16). The cocoa beans are ground/liquefied into a slurry (liquor) and heated, to ensure that all the butter is melted and a pumpable mass is produced. The liquor is pumped at a temperature of around 100 °C (212 °F) into the press until the press is fully charged. Pumping then stops and a hydraulic ram applies pressure to squeeze out the melted butter to 10–12% residual butter in the cake. Duyvis
Fig. 14.14 Insta-Pro extruder (courtesy Insta-Pro International).
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414 Separation, extraction and concentration processes
Fig. 14.15 Insta-Pro screw press (courtesy Insta-Pro International).
Fig. 14.16 Duvis press (courtesy Duvis Wiener BV).
has adapted this very much different type of press for pressing slurries of oilseeds other than cocoa bean, like peanuts and sesame seeds, with good results. The oilseeds are first ground to a slurry then heated to around 60 °C (140 °F). The heated slurry is then pumped into pots, and the pump pressure forces the vegetable oil to escape through fine-mesh grids built into the pots. The pump can exert a pressure of up to 2000 kPa (290 psi). Pumping then stops and a hydraulic ram forces the pots, with their contents, tightly together, to press out more oil. Final residual oils of 8–10% can be reached, but at © Woodhead Publishing Limited, 2010
Separation technologies in oilseed processing 415 the expense of capacity. Pressing cocoa to 22–24% residual oil can be done at 3290 kg h–1 (78 Mt d–1). Pressing to 10–12% residual oil can be done at capacities of 1315 kg h–1 (31 Mt d–1). The press is then opened, the solids are pushed into a conveyor, and the press closes for a new cycle. Cycling can be accomplished manually or automatically through a microprocessor.
14.5 Percolation solvent extraction in oilseed processing Solvent extraction is a very effective method for recovery of oil, especially with materials low in oil. Full-pressing can reduce oil content to a minimum of 3–5%. This is fine for oilseeds containing large amounts of oil, but for oilseeds low in oil (such as soybean at 18%), a reduction in oil from 18% to 3–5% leaves a reasonable amount of the available oil remaining in the presscake (approximately 16–27% of the total oil content). For that reason, solvent extraction is preferred for oilseeds of low oil content if capacities are sufficiently high to justify the expense of a solvent-extraction plant. Solvent extractors operate at lower temperature than full-presses, so the oil can be of higher quality if any components are heat sensitive. However, solvents do extract some components that are not true oils (triglycerides), thus necessitating some downstream procedures for removing them. Also, consumers are becoming reluctant to eat food products that have been in contact with chemical solvents. Percolation extractors work on the same principle as coffee percolators. Heated solvent is rained through the prepared oilseed, dissolving out the oil. Unlike a coffee percolator, however, the oilseed is extracted several times with the solvent redirected through the oilseed in a countercurrent flow pattern. Fresh solvent passes through the oilseed after it has been exposed to multiple passes of solvent. The fresh oilseed, on the other hand, as soon as it enters the extractor is washed with the solvent, now containing a reasonable amount of oil, just before the solvent/oil mixture (miscella) leaves the extractor. To ensure that the miscella contains a minimum of solids, most percolation extractors direct the final miscella one more time through the bed of solids (after the solids have had a chance to settle down into a firmer bed) in order to use the bed of solids as a filter to remove as many solid particles as possible. The extracted solids, after the final wash with incoming fresh hot solvent, remain in the extraction vessel to allow the solvent to free drain as much as possible before the finished extracted solids (marc) are removed from the extraction vessel. The marc then passes into a desolventization vessel. The miscella goes on to an evaporator followed by a stripper to remove all traces of solvent from the oil. Most percolation extractors permit external control of where the miscella from each stage goes. Valves downstream from each sump pump can send some miscella back to the stage from which it came and, by using different valves and different piping, send some to the preceding stage. Sometimes © Woodhead Publishing Limited, 2010
416 Separation, extraction and concentration processes all the miscella is pumped back over the stage from which it came. This allows the miscella to make multiple passes within each stage to increase solvent contact without requiring greater input of fresh solvent and also helps to maintain an adequate head of miscella above the bed of oilseed. The counter‑current movement of miscella from stage to stage, when the above is done, occurs through overflow weirs built into each sump. Oilseed preparation, especially particle‑size and moisture, influence how fast solvent can percolate through the bed of solids. Percolation should be rapid enough to ensure good contact of every particle with solvent. If percolation rate is too slow, the solvent will flood over the surface of the solids, greatly impairing the performance of the extractor. Similarly, if percolation is so rapid that the solvent flows down through the center of the bed without spreading across the entire bed, some of the oilseed will not come into contact with the solvent. Marc from flaked oilseeds usually retains 30 to 40% by weight of solvent. Pre-pressed cake and extrusion prepared oilseeds can free drain to 20–25% solvent before leaving the extraction vessel. This lower retention of solvent by the marc is one of the major benefits of extrusion before solvent extraction. If the marc free drains to a low solvent level, substantial energy can be saved in desolventization because less solvent has to be driven off by thermal evaporation. However, the solvent that drains out of the marc ends up in the miscella, which increases the solvent that has to be evaporated from the miscella. In all instances, where extrusion preparation is used, the solvent flow into the extractor can be lowered, resulting in less solvent in the miscella as well as less solvent in the marc. Extruded collets permit this reduction in solvent-to-meal ratio because extrusion helps release the oil from the solids, enabling most of the oil to extract very quickly. Rapid percolation also helps by flushing out the released oil almost as soon as the solids enter the extractor. The more deeply embedded oil still needs time to be fully extracted, but less solvent is required, thus permitting an extraction plant to reduce energy consumption for both miscella and marc desolventization. 14.5.1 Rotary extractors Rotary extractors have a rotating circular disk-shaped bed of oilseed slowly spinning within the extraction chamber. These extractors are quite large, and the bed depth of oilseed can be 2–4 m. Fresh hot solvent is directed over the bed at the end of the extraction cycle, just before the marc is allowed to free drain. The oilseed is extracted in a counter-current fashion as described above. Figure 14.17 shows a schematic of a typical rotary extractor. 14.5.2 Perforated belt extractor Another type of percolation extractor uses a longitudinal, perforated belt carrying an extended bed of oilseed. The belt forms a flexible endless loop
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Separation technologies in oilseed processing 417 Top and bottom bearings accessible from outside the unit for easy maintenance
Slurry filling spout
Ultra-free turning spindle
Sealed dividers form baskets to ensure miscella stage separation
Conical top with sight glasses for maximum visibility
Reliable bevel gear drive Self-cleaning screen for outstanding drainage Miscella collection pan
Sealed dump hopper to prevent contamination
Fig. 14.17 Schematic of rotary extractor (courtesy Desmet Ballestra).
moving over pulleys at both ends and carrying the oilseed on the top surface through the extraction vessel. The belt is on a slight incline, and the solids travel up the incline toward the discharge end. This helps to ensure a more efficient countercurrent flow of solids and miscella while in the extractor. The marc falls off the higher end as the belt passes around the discharge pulley. Fresh oilseed is dropped onto the lower end of the belt after it comes to the top of the feed end pulley. 14.5.3 Sliding cell extractor A sliding cell extractor slides the solids along a stationary steel plate that is perforated to allow miscella to drain through the plate while the bed of solids
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418 Separation, extraction and concentration processes slide across the surface of the plate. The incoming oilseed falls onto an endless belt mechanism consisting of tall vertical plates that form closed receptacles or cells into which the oilseed is captured. Countercurrent miscella washes are introduced onto the bed of cells as they pass under miscella spray heads. The moving cell assembly wraps around a drive pulley, turning the oilseed contents upside-down and sliding them across a second perforated plate as the cells return back to the feed end of the extractor. Each cell empties just before the moving assembly wraps around a second pulley. The now empty cells then move under a hopper from which fresh oilseed enters the cells. 14.5.4 Rectangular loop extractor A rectangular loop extractor drags the oilseed through a chamber shaped like a closed loop. This extractor uses an ‘en masse’ conveyor that moves the oilseed through a closed housing with a rectangular cross section. The housing loops back on itself, looking somewhat like a collapsed doughnut resting on its side. Fresh oilseed is conveyed through an inlet onto the upper level and is sprayed with rich miscella. The oilseed encounters multistage counter-current extraction as it travels through the loop. The final wash is with fresh solvent followed by a free drain period before the extracted oilseed exits. Figure 14.18 shows a rectangular loop extractor. 14.5.5 Extractor suppliers Major suppliers of solvent extraction equipment are Crown and Desmet. Other suppliers are Harburg-Freudenberger and Lurgi GmbH and several others around the world. A list of suppliers of oilseed processing equipment is available (Anon, 2009b). Many of these companies also supply auxiliary equipment and can supply entire plants. Crown specializes in rectangular loop extractors (Fig. 14.18). Crown offers their Model 3 extractor, which is rated for 9000–9500 Mt d–1 on flaked oilseeds and up to 10 000 Mt d–1 on extruded collets. Crown has made advances in hexane recovery, especially in stripping and desolventation. Their Super Stripper System™ can result in significantly lower solvent losses than the 3.78 L t–1 (1 gal ton–1) that was common 10 years ago. If the equipment is operated correctly and kept in good repair, solvent losses of 0.8–1.9 L Mt–1 (0.22–0.5 gal ton–1) can be achieved. Desmet makes two types of solvent extraction vessels: the Reflex® solvent extractor, a rotary extractor, and the LM™ extractor, a perforated belt extractor. Capacities of Desmet extractors range from 500 to 10 000 Mt d–1 on flaked soybean. Desmet’s Reflex® rotary extractor has a horizontal circular extraction chamber divided into sections like a cut pie. The material to be extracted revolves under a stationary solvent/miscella spray harness and feed input chute, which directs the incoming oilseed and solvent/miscella over the appropriate sections of oilseed as the bed of material rotates under
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Solids inlet hopper with electronic level sensor
Extractor drive speed controlled by the incoming volume of raw solids read by the electronic sensor
First wash
Hydroclone
Hydroclone miscella clarifier Full miscella outlet
Countercurrent recycle stages
The Crown Hydroclone removes the final traces of fines from the full miscella. The miscella can be pumped directly to the evaporation system.
Fresh solvent rinse
Drainage section
Final recycle
Bar screen Self-cleaning Countercurrent stationary vee-bar recycle stages screens
The flake bed acts as a brush - it continually shears the stationary bar screen clean of flow-obstructing fines
Fig. 14.18 Rectangular loop schematic (courtesy Crown Iron Works Company).
Separation technologies in oilseed processing 419
© Woodhead Publishing Limited, 2010 Extracted solids outlet
420 Separation, extraction and concentration processes the spray harness. The oilseed within the pie-shaped sections slides over a stationary perforated bottom. The sliding of the oilseed over the bottom helps to keep the perforations open. When extraction time is finished, the extracted oilseed (marc) slides over a lower stationary marc hopper, and the marc drops into the hopper. The rotating extraction chamber then moves under the stationary feed inlet hopper, which drops fresh oilseed onto the revolving disk. Desmet also supplies perforated belt extractors. The belt is made of a series of wide, rectangular sections of either wire mesh or wedge-bar screens that are connected together to form a looped endless-conveyor belt. Figure 14.19 shows the Desmet LM™ extractor, and Fig. 14.20 shows a schematic of this extractor. The perforated belt extractor has only one moving part: the belt. Belt speed is automatically adjusted to maintain a constant level in the inlet hopper. The belt is self-tensioning, self-cleaning, and fully automatic. The extractor is designed to permit various extraction and drainage times before the marc discharges from the extractor. Typical times for soybean are 42 min for extraction and 15 min for drainage. Other duration times are set for different oilseeds. Harburg offers their carousel extractor, which is a rotary extractor with the chamber containing the oilseed rotating beneath a stationary spray head/input hopper assembly and above stationary miscella sumps and marc hopper assembly. A carousel extractor can be supplied double-decked, and its capacity is 50 to 4000 Mt d–1. Lurgi offers a line of sliding cell extractors with a capacity range of 100
Fig. 14.19 Perforated belt extractor (courtesy Desmet Ballestra).
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1
2
3
2
Fig. 14.20 Perforated belt schematic (courtesy Desmet Ballestra). 1, Fresh oilseed input; 2, fresh solvent input; 3, miscella output; 4, marc output.
Separation technologies in oilseed processing 421
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422 Separation, extraction and concentration processes to 4000 Mt d–1. The Lurgi sliding cell extractor provides dual-pass, shallowbed extraction with the partially extracted oilseed, at the midpoint of the extraction cycle, refilling the cells in order to ensure better solvent contact with the oilseed. The cells containing the oilseed slide over rod-type screen plates with conical rods to minimize clogging of the screen plates. 14.5.6 New developments in solvent extractors Newer developments in solvent extraction plants are geared to very large extractors permitting 6000 to 10 000 Mt d–1 capacity on soybean and very efficient recovery of solvent.
14.6 Solvent recovery in oilseed processing 14.6.1 Recovery from miscella Hexane is the most frequently used solvent to extract oils. Other solvents have been considered (Hron et al., 1982; Johnson and Lusas, 1983). First step in recovering solvent from the miscella is to remove solid particles (fines). This can be done before or after solvent removal. However, if fines are left in the miscella and only removed after desolventization, they cause problems in evaporators and strippers, and fines remaining after desolventization contain more oil (that will be subsequently lost) than fines removed from the miscella. Solvent is usually removed from the miscella in a two-stage rising-film evaporator, followed by a ‘packed’ or a ‘disk and doughnut’ stripping column, which remove the last traces of solvent. The first-stage evaporator can be heated by hot solvent vapors from the marc desolventizer. These hot solvent vapors can also be used to preheat the miscella in a heat exchanger. The second-stage evaporator is heated by jacket steam. The solvent vapors from both evaporators are condensed whereas the oil from the evaporators, containing about 5% solvent, goes into the stripping column and is exposed to a counter-current flow of live steam to strip out the last traces of solvent. In the stripping column, the oil forms a thin film, giving it a large surface area as it flows over the packing, facilitating the stripping. Stripping columns are operated under high vacuum, 559–711 mm (22 to 28≤) Hg to help vaporize the solvent. Solvent and steam vapors from the stripper are condensed and pass to a ‘solvent/water separator’. The condensed liquids separate into two phases: hexane, being less dense, floating above the water. As the separator fills, excess hexane overflows from the top and goes to the solvent work tank. The separator height is chosen so that, at a predetermined liquid level, the weight of the liquid will force water at the bottom to pass up a siphon tube and go to a water stripper. There are still traces of hexane in the water from the solvent/water separator, so this water is scrubbed with live steam to ensure that no traces of hexane remain in water leaving the solvent plant. © Woodhead Publishing Limited, 2010
Separation technologies in oilseed processing 423 The steam from the water stripper is condensed and sent to the solvent/water separator tank. 14.6.2 Recovery from marc The marc is transported to a ‘desolventizer/toaster’, which is divided into several horizontal sections or ‘trays’ within a vertical cylindrical vessel. Sweep arms, attached to a vertical shaft skimming across the surface of each tray, keeps the marc agitated. The sweep arms also trip trap doors in each tray to allow marc to fall into the lower trays. Tray bottoms are steam jacketed and sometimes the vessel side walls also. Sparge steam, introduced into the top stack, helps to flash off the solvent. The solids absorb the condensed sparge steam, minimizing dust carryover to the condenser. The sparge steam also cooks or ‘toasts’ the solids. This helps to inactivate enzymes, like urease in soybean. Some desolventizers use a counter-current flow of live steam, introduced under the bottom tray and passing upward through perforations in each tray and through the marc on each tray. This greatly assists desolventization. The perforations are through hollow staybolts dispersed into each tray. The top tray has a steam-jacketed bottom. The heat emanating from the jacket helps flash off surface solvent in the incoming marc. Another design, a flash desolventizer, vaporizes solvent at low temperature to prevent any denaturing of the protein. The marc enters a recirculating stream of hot solvent vapor. Most of the solvent in the marc vaporizes. Some of the vapor, equal to what is flashing from the marc, is bled off through a rotary valve. These solvent vapors are then scrubbed to remove fines and liquefied in a condenser. The marc solids, still at about 1% solvent, are directed through a cyclone and rotary valve, then pass into a ‘flake stripper’ where the final traces of solvent are removed under low-heat, low-moisture conditions that will not denature or darken the protein. 14.6.3 Recovery from effluent air Air enters the solvent extraction system with the incoming material and through any leaks in the vessels, many of which are operating under a vacuum of about 2.4 to 5 mm (1 to 2≤) water column. This ensures that, if there are leaks, air will leak into the equipment rather than solvent vapors leaking out. This air will leave the solvent extraction system saturated with solvent vapor. Effluent air collected from all vessels likely to contain incoming air, such as the extraction vessel and the desolventizer, is directed into a common header, and then passes through a device designed to remove as much solvent as possible before the air is discharged. Two devices are preferred for maximum solvent recovery. One device is like a miniature solvent extraction system using mineral
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424 Separation, extraction and concentration processes oil as a solvent to absorb, and thus extract, hexane from the effluent air. The mineral oil is then stripped in a similar way to how vegetable oil is stripped. The stripping steam and hexane vapor are then condensed and pass to the main solvent/water separator. The hot mineral oil from the stripper is cooled and returns to the mineral oil scrubber. In this fashion, the mineral oil is continuously recirculated through the absorber, the stripper, and the cooler. The second device passes the vented air, still containing solvent vapor, through a bed of activated charcoal. Most of the solvent is absorbed by the charcoal while the air passes through. When the charcoal is almost saturated with solvent (and no longer effective), the incoming air is diverted to a second charcoal absorber, which is piped in parallel with the first one. The air from both absorbers is scrubbed with sparge steam to strip out the solvent. That solvent/water mixture is then condensed and sent to the main solvent/water separator. The absorber saturated with solvent undergoes a cleansing cycle to remove the absorbed solvent. Using either device, solvent loss from the processed oilseed through effluent air can be kept below 1.89 L t–1 (0.5 gal t–1), and, if the equipment is properly maintained and operated, solvent losses as low as 0.8–1.89 L Mt–1 (0.22–0.5 gal t–1) can be achieved. 14.6.4 Recovery from effluent water All mixtures of steam and solvent vapors, after being condensed, are sent to a common solvent/water separator where solvent overflows the top, and water siphons out the bottom, as described above. The water is then passed through a steam-heated ‘waste water stripper’ where live steam elevates the water temperature above the boiling point of the solvent but under the boiling point of water. The desolventized water then leaves the solvent extraction system.
14.7 Obtaining oil from fruit pulps Typical fruit pulp oils are olive oil and palm oil. These oils reside in the soft fruit pulp rather than in the kernel or seed. Palm fruit oil serves the same markets as other vegetable oils. Olive oil is valued for its flavor. 14.7.1 Palm oil Palm fruits grow in clusters on a central stalk. The fruit is composed of oily pulp within a tough outer skin and has seeds, or kernels, imbedded in the pulp. Two very different types of vegetable oil are obtained from this palm plant. Palm fruit oil comes from the pulp. Palm kernel oil comes from the seed. © Woodhead Publishing Limited, 2010
Separation technologies in oilseed processing 425 The clusters of freshly picked palm fruit are sterilized in steam chambers and then are sent through the stripping/threshing equipment to remove the individual fruits. The separated fruits are next washed to remove sand. Pressing is done in a twin-screw press to liberate the palm fruit oil. The twin screw press is operated at lower pressure than full-presses. The final step is to clarify the oil of moisture and debris. The twin-screw press is equipped with two parallel worm shafts that rotate in opposite directions and transport the fruit through drainage cages that surround the shafts. Hydraulically operated cone chokes apply back pressure. The liquid, a mixture of palm fruit oil and water, flows out through the drainage cage. Solid residue is pushed over the cone chokes by the action of the shafts. The expressed mixture of oil, water, and some solid impurities, is further processed to clarify the oil. The solids from the press consist of moist pulp solids, the outer skin, and the palm kernels. The solids are sent to a pneumatic separator to separate the kernels from the fiber and other debris. The kernels are cracked to separate the meats from the shells. The meats, still moist, are dried to 7% moisture and stored for subsequent full-pressing. 14.7.2 Olive oil Previous methods involved washing the olives, crushing them, stirring them into a thick paste, and then pressing to separate olive oil from the paste. Pressing was done in batch-operated plate presses by inserting bags between the plates followed by a hydraulic ram to squeeze out the oil. The oil was then centrifuged and clarified using diatomaceous earth in a subsequent filter press. Currently, centrifuges are used. The olives are first washed and then pass through a milling and beating stage followed by centrifuging. In the 1970s, three-phase centrifuges were used, employing hot water to obtain maximum oil yield. The wastewater stream presented ecological difficulties so, in the 1990s, the three-phase centrifuges were replaced by two-phase centrifuges, which greatly reduced the amount of wastewater (Harmsen and Mulder (2009, p. 20).
14.8 Future trends Recently, new objectives and new concerns have influenced oilseed processing, and equipment suppliers are developing new equipment and new procedures to address these concerns. Currently, most vegetable oils serving the food market are separated by solvent extraction because solvent extraction permits large capacity plants and maximum separation of oil. However, recent concerns about foods exposed to hexane and other chemicals, especially chemicals known to have deleterious consequences to the consumer, have © Woodhead Publishing Limited, 2010
426 Separation, extraction and concentration processes caused consumers to become suspicious of foods exposed to chemicals. Today, many foods are labeled ‘organic’ signifying that the foods have not been exposed to chemicals such as fertilizers, pesticides, solvents, and some preserving chemicals widely used in the past. This trend will probably continue, encouraging food producers to prepare foods without exposure to any chemicals. For that reason, future processors might favor screw pressing over solvent extraction (if screw pressing could be done at higher capacities and at lower cost) because no hexane or other chemical solvent is introduced into the food material being processed. Screw presses have traditionally been low-capacity machines compared with solvent extractors and consume more horsepower per tonne of oil extracted than solvent extractors. There is an incentive, today, for screw press manufacturers to develop higher capacity screw presses that consume less horsepower. Some manufacturers are already working on this. Other manufacturers are developing novel procedures, such as injecting liquid supercritical carbon dioxide through the oilseed within the screw press to maximize release of oil. This and other new developments have been discussed in this chapter. A new market is developing for the use of vegetable oil to replace petroleum products so that bio-renewable fuels (oils for diesel engines and ethanol from fermentation of the solid residue for gasoline engines) can diminish world dependence on fossil fuels. There has been some success producing ethanol from cereal grains such as maize and diesel oils from sunflower and rape seed. Success in these areas led to large-scale plants for making ethanol from maize starch and many small screw press plants to obtain vegetable oils for biodiesel. Some of the more efficient biodiesel plants also contain esterification equipment. Success in obtaining bio-renewable fuels from traditional food crops shows signs of placing greater demand on available food crops thereby raising the price for these food crops. This could have unfortunate long-term results. Also, the fossil fuel market is huge and could eventually consume a very large portion of available food crops. Therefore, efforts are currently being made to identify oil-containing materials that have not been processed before and could easily serve the bio-renewable fuel market without compromising the edible food market. A very promising source of untapped biorenewable oil is from algae. Algae are easy to grow, have sufficient oil content, and can, with some difficulty, be processed to obtain the oil. The world’s first algae biodiesel plant went online in Rio Hondo, Texas on April 1, 2008 (Cornell, 2008). It will make ethanol from the algae biomass first. Future plans involve also separating the oil. Algae has potential because, if and when proven successful, algae grown on only 9.5 million acres of land could replace all transportation fuels consumed annually in the USA. Farm cropping in the USA, in contrast, currently uses 450 million acres (Briggs, 2004). Recently, the New York Times reported that Exxon Mobil Oil Company, once skeptical of algae, has announced a 600 million US dollar program to produce biofuels from algae (Anon., 2009b).
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Separation technologies in oilseed processing 427
14.9 Sources of further information and advice Much has been written about standard equipment and procedures. There is no need to describe once again what is already well known and fully described in earlier books. For those interested in what was done in the past, The recovery of oils and fats from oilseeds and fatty materials (Shahidi, 2005) contains a thorough description of the early development of screw presses and solvent extractors and the procedures used from early development to state-of-the-art practiced at time of publication. Another good reference is Gunstone and Padley, (1997). Other good references are Modern technology of oils, fats & its derivatives, (NIIR, 2002) and Technologies for the recovery of residual oil (Harmsen and Mulder, 2009). Extrusion as preparation ahead of screw pressing and solvent extraction can also be further investigated through Farnsworth et al. (1986), and Lusas and Watkins (1988); Nelson et al. (1987), Williams (1993). Hexane is the most widely used solvent for solvent extraction of vegetable oils. Information about other solvents that are used can be found in Shahidi (2005) and Hron et al. (1982). Some manufacturers of screw presses, solvent extractors, and preparation equipment are identified in this chapter. Others can be found in Anon. (2009a) and on the Internet. Suppliers of plant and equipment Anderson International Corp, 6200 Harvard Avenue, Cleveland, OH 44105, USA (http://www.andersonintl.net). Crown Iron Works Company, P.O. Box 1364, 1600 Broadway Street N.E., Minneapolis, Minnesota 55440-1364, USA (http://www.crowniron. com). Desmet Ballestra Group N.V. Minervastraat 1, B – 1930 Zavetem, Belgium (http://www.desmetballestra.com). The Dupps Company, 548 N. Cherry Street, Germantown, OH 45327, USA (http://www.dupps.com). Duyvis, B. V. Machinefabriek P. M. Duyvis, 1541 KD Koog ann de Zaan, P.O. Box 10, The Netherlands (http://www.pmduyvis.nl). The French Oil Mill Machinery Co., 1035 West Green Street, P.O. Box 920, Piqua, Ohio 45356-0920, USA (http://www.frenchoil.com). Harburg Krupp Maschinentechnik GmbH, Werk Harburg, Postfach 900880, Seevestrasse 1, D-2100 Hamburg 90, Germany (http://www.harburgfreudenberger.com). Insta-Pro International, 10104 Douglas Avenue, Des Moines, IA 50322, USA (http://www.insta-pro.com). Lurgi GmbH, Lurgiallee 5, ED-60259 Frankfurt am Main, Germany (www. lurgi.com).
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14.10 References Anon. (2009a) NY Times Available from: http://www.nytimes.com/2009/07/14/business/ energy-environment/14fuel.html (Accessed 15 July 2009). Anon. (2009b), ‘Plant and equipment survey’, Oils & Fats International, 25, 38–41. Baer S, Williams M and Zies C (1966), ‘Pre-treatment of oleaginous plant materials’, US Patent 3,255,220. Briggs M (2004), Wide scale biodiesel production from algae, University of New Hampshire, Physics Dept. Available from: http://www.unh.edu/p2/biodiesel/article_alge.html (revised August 2004) [Accessed 16 July 2009, note algae misspelled] Cornell C (2008), First algae biodiesel plant goes online: April 1, 2008, Available from: http://gas2.org/2008/03/29/first-algae-biodiesel-plant-goes-online-april-1-2008/ (Accessed 16 July 2009). Farnsworth J, Johnson L, Wagner J, Watkins L and Lusas E (1986), ‘Enhancing direct solvent extraction of oilseeds by extrusion preparation’, Oil Mill Gaz, 91, 30. Friedrich J, List G and Heakin A (1982), ‘Petroleum-free Extraction of Oil from Soybeans with Supercritical CO2’, J Am Oil Chem Soc, 59, 288. Friedrich J and Pryde E (1984), ‘Supercritical CO2 Extraction of Lipid-bearing Materials and Characteristics of the Products’, J Am Oil Chem Soc, 61, 223. Gunstone F and Padley F (1997), ‘Extraction of lipids from natural sources’, Lipid Technologies and applications, New York, Marcel Dekker, Inc., 113–135. Harmsen P and Mulder W (2009), Technologies for the recovery of residual oil, Available from: http://www.york.ac.uk/res/sustoil/Pages/Deliverable%202.3%20FINA.pdf (Accessed 16 July 2009). Hron R, Koltun S and Graci A (1982), ‘Biorenewable solvents for vegetable oil extraction’, J Am Oil Chem. Soc 59, 674A. Johnson L and Lusas E (1983), ‘Comparison of alternative solvents for oils extraction’, J Am Oil Chem Soc 60, 229. Lusas E and Watkins L (1988), ‘Oilseeds: Extrusion for Solvent Extraction’, J. Am. Oil Chemists’ Soc. 65, 1109. Nelson A, Wijeratne I, Yeh W, Wei T and Wei L (1987), ‘Dry extrusion as an aid to mechanical expelling of oil from soybeans’, J. Am. Oil Chemists’ Soc. 64, 1341. NIIR (2002), Modern technology of oils, fats & its derivatives, New Delhi, National Institute of Industrial Research, India, ISBN: 8178330857. Radio Iowa, ‘Elsworth company breaks ground for soy plant,’ Available from: http://www. radioiowa.com/gestalt/go.cfm?objectid=57F7E7F2-B31D-4124-ADC4E78B4C8934B1 (Accessed 16 July 2009). Shahidi F (2005), ‘The recovery of oils and fats from oilseeds and fatty materials’, Bailey’s industrial oil and fat products, Sixth Edition. Edited by Fereidoon Shahidi, New York, John Wiley & Sons, Vol. 6, 2589–2678. Vadke V, Solulski F and Shook C (1988), ‘Mathematical simulation of an oilseed press’, JAOCS, 65, 1610. Also available from: http://www.springerlink.com/content/ b71x7j050272mvjn/ (Accessed 16 July 2009). Voges S, Eggers R and Pietsch A (2008), ‘Gas assisted oilseed pressing’, Sep Purif Technol, 63, 1–14. Willems P (2007), Gas assisted mechanical expression of oilseeds, Universiteit Twente, Nederland. Available from: http://doc.utwente.nl/58041/(Accessed 16 July 2009). Williams M and Baer S (1965), ‘The expansion and extraction of rice bran’, J. Am. Oil Chemists’ Soc. 42, 151. Williams M (1983), ‘Description of Anderson International’s vector shaft analysis technique’, Oil Mill Gaz, March, 36–37.
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Separation technologies in oilseed processing 429 Williams M (1990), ‘Apparatus and method for the continuous extrusion and partial deliquefaction of oleaginous materials’, to Anderson International Corp. US Patent 4,901,635. Williams M (1991), ‘Extruded starter pig feeds’, Feed Manage 42, 20. Williams M (1993), ‘Preparation of oilseeds to improve extraction of fats’, Extrusion Commun 6, 12.
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15 Separation technologies in brewing G. J. Freeman, Campden BRI, UK
Abstract: The series of separation processes that characterises the brewing process is overviewed with emphasis on those processes that provide opportunities for technological progress. Extraction of raw materials in the brewhouse and the filtration and inertial separation processes employed are described. Yeast separation after fermentation is mainly achieved through natural flocculation processes, which may be enhanced by fining agents. Filtration to commercial quality beer clarity may be achieved by filter aid filtration, including novel filter aids that are more environmentally friendly, or membrane filtration. Other applications of membranes include control of dissolved gas levels and recovery and reuse of detergents. Key words: wort separation, yeast flocculation, beer filtration, membrane technology.
15.1 Introduction The brewing process is characterised by a sequence of many separation processes. The processes of brewing (hot extraction), fermentation, maturation and end processing generally comprise two or three phases of solid, liquid or gas. Processing requires accurate control of the appropriate variables in order to achieve the desired quality and consistency in final product. Sales of beer globally have become dominated by stable products with long shelf life, perhaps as long as a year. This requires a high standard of hygiene, process equipment and operating procedures. Shelf life is limited by either microbiological, flavour or colloidal (turbidity) instability. Brewers are fortunate in that no pathogens can grow in mainstream beer products. However, there are a variety of micro-organisms that can spoil beer by causing
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Separation technologies in brewing 431 turbidity, acidity and other off-flavours (O’Rourke, 2000). Flavour instability is a complex deterioration of the product by a variety of chemical reaction schemes. ‘Stale’ beer is characterised by the loss of desirable bitterness and a variety of undesirable flavours such as ‘wet cardboard’ and ‘winey’ (Bamforth, 2000). Colloidal instability may occur in a variety of ways, but the most important, and that which practically all beer is susceptible to, is the production of particles by flocculative interactions between polyphenols and certain polypeptides (Siebert, 1997). The microbiological, flavour and colloidal stability is extremely adversely affected by the presence of oxygen in the final product. There is a need to exclude oxygen as much as possible after the process of primary fermentation. Final packaged beer should if possible have as little as less than 0.1ppm of oxygen. Again, this necessitates rigorous process plant design and procedures. Most beer is pasteurised, and this, along with hygienic plant design and operation, normally achieves microbiological stability over the shelf life of the product. Flavour stability may perhaps be quantified by the measurement of concentrations of ‘marker’ chemical compounds (Malfliet et al., 2008) but this is complex and still under development. Therefore, colloidal stability is usually employed as the definitive measure for product shelf life. There is a need to process the polyphenols and polypeptides that are responsible for colloidal instability by allowing the formation of the particles and then removing them by a variety of separation process options. Indeed, the more effective the separations throughout the process, even very early in the process, the better the colloidal stability of the beer.
15.2 Characteristics of brewery products There are a large number of distinctive beer styles. They are characterised not just by region of origin, but by differences in process and packaging style that are particularly relevant in this chapter. Most beer in the world is retailed in small packs of glass bottles, stainless-steel or aluminium cans or plastic (most commonly polyethylene terephthalate) bottles. These products are normally filtered to an appealing visual clarity. Cans and glass bottles are most commonly pasteurised in package (tunnel pasteurisation). This is not always possible with plastic bottles, although in recent years some bottles amenable to high-temperature processing have come on to the market (Martin, 2002). Plastic bottles are most commonly sterile filled after sterile filtration. Larger containers usually employed in bars and restaurants are varieties of keg. Microbiological stabilisation is not possible in these containers, but is achieved by sterile filtration or, more commonly, by in-process pasteurisation (flash pasteurisation by plate heat exchanger). In-process pasteurisation is more thermally efficient than in-pack. Thus, there are drivers to employ flash pasteurisation in these times of increasing energy prices (Browne, 2008). © Woodhead Publishing Limited, 2010
432 Separation, extraction and concentration processes Some beers, normally originating from discrete regions, are characterised by not having been filtered. Cask-conditioned beers undergo secondary fermentation (maturation) in the container. The yeast is sedimented from the beer by employing flocculents known as finings. Bottle-conditioned beers undergo secondary fermentation in the bottle.
15.3 Selection of technology and raw materials appropriate to brewery products The main raw material employed in beer production is malted barley. The malting process involves initiating the growth of the barley plant induced by increasing the content of water by steeping. The grain is allowed to grow (germination) for a period of some days, which causes structural and enzymic changes in the grain that enable the malt to be processed in the brewhouse. These changes, known as modification, are allowed to occur up to a certain point dependent on the specification required for the beer product. Finally, the process is stopped by drying with hot air (kilning) that produces a stable product. The time–temperature cycle employed in kilning also controls aspects of the malt character notably flavour and colour. Malts that are employed to make darker beers such as ales are typically more modified and kilned to higher temperatures. In the brewery, the malt is milled with either roller mills or hammer mills. The detailed choice of mill depends upon the extent of modification of the malt and the technology employed in the brewhouse. The milled malt, known as grist, is contacted with warm water in a process known as mashing. The grist may be supplemented with other processed cereals such as wheat or maize. In the mash, several biochemical reactions occur as a result of the enzymes in the grist. The result is a sugar- and nutrient-rich medium known as wort. The wort is separated from the remaining undissolved brewers’ grains by filtration through the bed of grains itself. The technology selection is limited by the choice of raw materials as described herein. The wort must undergo a boiling process. This is required for reasons of flavour, production of bitterness and flavour from hops that are traditionally added to the boil and for stability and sterilisation reasons. Sugar-rich syrups may also be added to supplement the wort. Excess protein precipitates as large loose solids known as trub that requires removal along with the remaining hop solids. This separation process is most commonly performed in a vessel known as a whirlpool described herein. The wort is then cooled, transferred into the fermenter vessel and pitched with active yeast. The fermentation process is exothermic and is controlled by a temperature programme. Some darker beers such as ales and stouts may be allowed to reach temperatures greater than 20 °C, whereas typical lagers are not fermented above 15 °C. Thus, the lager fermentation takes longer
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Separation technologies in brewing 433 but produces a more delicate flavour, an effect that is exaggerated by the employment of less kilned, less flavoursome malt products. Separation of most of the yeast occurs through the natural flocculation and either sedimentation or flotation of the resultant flocs. A secondary fermentation is required for reasons of flavour. Most beer then undergoes a cold stabilisation (storage) period. This is required to precipitate more excess protein by flocculation with polymerised polyphenols. This is required for products that are served chilled (most beer) and/or require extended shelf-life. Most beer is clarified by fine filtration. Some employ flocculents in package, known as finings, that maintain suspended solids in a large floc away from the dispense point. Cask-conditioned ales are the main example of this. Bottle-conditioned beers are similar, also containing live yeast as a sediment but without finings. Microbiological stabilisation is required if yeast is not present in package to inhibit micro-organism growth. This occurs by pasteurisation or sterile filtration before packaging.
15.4 Wort production in the brewhouse After the biochemical reactions have occurred in the mashing process, the liquid wort is separated from the brewer’s grains by filtration through the grain bed itself. The majority of the extract is recovered to the wort by rinsing (sparging). The filtration rate, which commonly limits the capacity of the brewhouse, is mainly controlled by both the coarseness of grind and the depth of the bed. The three dominant wort separation technologies are mash tuns, lauter tuns and mash filters. The relative merits of the three processes have been reviewed by O’Rourke (1999) and are described hereafter only in broad classification terms. A mash tun is characterised by mashing and wort separation occurring in the same vessel. The grain and hot water are mixed in such a way that air is entrained and the grain bed floats. Sparging occurs by spraying hot water on top of the bed. The mash tun is limited to highlymodified, coarsely ground malt (a fraction may be from other cereal starch sources) that makes it most suitable for the production of ales. The coarse grind gives a thick mash and a deep filter bed (1 m deep). A lauter tun involves similar technology to a mash tun, but the mashing process proceeds in a separate vessel called a mash mixer before pumping into the lauter tun. The lauter tun is more flexible on the malt qualities that may be employed and, for example, is more appropriate to many lagers that are made from less well modified malts. The grind is less coarse than for a mash tun and hence the bed must be shallower (40 cm). Flexibility in choice of raw materials is further improved by the employment of raking of the bed if the filtration rate is found to be too slow. Mash filter technology has been greatly advanced over the last three decades. The grain is ground to a relatively fine flour by a hammer mill (most commonly) as opposed to roller mills employed for © Woodhead Publishing Limited, 2010
434 Separation, extraction and concentration processes mash tuns and lauter tuns. The fine grind necessitates the need for a thin filter bed. The advances have mostly centred around pressurisation of the filter bed to increase filtration rate by, for example, pneumatically driven polymeric sheets (Eyben et al., 1989). The lauter tun and mash filter technologies allow much greater throughput than mash tuns. Mash filters have been proposed as part of a continous brewing technology, being operated semi-continuously with one in forward flow with the other being readied for use (Harmegnies et al., 2003).
15.5 Whirlpools and applications in brewing There is a need to remove solid material after boiling the wort in the brewhouse. Excess protein is removed as precipitated trub, which if it remains through to fermentation will cause sulfurous off-flavours and complicate downstream processing. Most breweries employ whirlpools, the mechanism of which is described in Fig. 15.1. The vessel is filled through a tangential main and then allowed to stand, the liquid maintaining a rotational flow, for a period typically of 15–60 min. Clear wort may then be drawn off away from the accumulated solids at the bottom centre of the vessel. The trub solids exist as loose, delicate flocs. With all flocculation processes, it is necessary to ensure that the correct shear forces are applied in the transferring pipe mains and pumps. Excess mixing results in small flocs that are difficult to separate, typically manifesting as a poor ‘cone’ of
Continuing liquid rotation after vessel is full causes concave liquid surface
Resultant liquid flow. Upflow at centre insufficient to lift accumulated solids
Tangential inlet(s) for fill.
Force due to larger liquid head nearer side of vessel
Force due to particle density being higher than wort density
Force balance causes solids accumulation bottom centre
Fig. 15.1 Mode of operation of a whirlpool for trub separation.
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Separation technologies in brewing 435 accumulated solids smeared across the base of the vessel. Inadequate mixing means that the trub particles do not contact frequently enough, resulting in similar separation problems (Virden, 1995).
15.6 Yeast flocculation and applications in brewing The traditional brewing process relies on the timely flocculation of yeast in order to clarify the beer sufficiently so that downstream processes are able to finish the beer to consumer satisfaction. Thus, it has always been the case that a major selection criterion for a suitable brewing yeast is that it flocculates at the correct stage of the fermentation (Soares, 2009). The application of modern centrifuges means that latterly it has been possible to employ less flocculent yeast strains. The flocculation process itself comprises interaction between polysaccharides and glycoproteins on the cell surface (Evans and Kaur, 2009). The timing of rapid flocculation is controlled by factors such as the presence of calcium ions, a low level of unfermented extract and temperature. Some relatively hydrophobic strains, such as those that are traditionally employed to make ales, may be induced to flocculate to the surface of the wort to be removed by ‘skimming’. Most commonly, brewing yeast strains flocculate to the bottom of the vessel. It is in the nature of all flocculation processes that they are affected by mixing in the system. The flocculent particles require a degree of mixing so that they contact and join together to form flocs. Mixing is provided in the fermentation vessel by the evolution of carbon dioxide. However, it has been suggested that a degree of forced mixing is beneficial (Boulton, 2009). Pumping the fermenting wort around a loop around the vessel appears to result in: more efficient cropping of the yeast (presumably making downstream processing less expensive also), reduced fermentation cycle times, increased ethanol yield and the ability to cool the fermentation more rapidly at the end of the cycle because of improved heat transfer. One interesting phenomenon that is currently under intense investigation is premature yeast flocculation (PYF). This effect has probably always intermittently appeared without being correctly diagnosed. The consequences are serious, with the yeast unable to complete fermentation resulting in excessive residual carbohydrate and other undesirable beer flavours (Lake and Speers, 2008). At the time of writing it is still not well understood. Most likely, the problem originates from microbial, probably fungal, contamination of barley in the field (Evans and Kaur, 2009). The PYF-inducing factor, or possibly more than one factor, appears to have molecular weight less than 100 kDa. However, the effect of process factors in both malting and brewing are unclear. In the germination process during malting, air temperature, air flowrate and humidity have been implicated as conditions that affect the presence of the PYF factor. However, malts that induce PYF in some brewing
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436 Separation, extraction and concentration processes processes may not in others. There are probably differences caused by the final wort quality and the conditions in the fermentation vessel.
15.7 Beer fining agents Finings are flocculents in the brewing process. The most important is isinglass. This is a remarkable substance that is manufactured by solubilising the swim bladders of fish into dilute mineral acid. The result is a viscous suspension of almost pure collagen. When added to beer, the collagen molecules adopt a net positive charge. The large majority of particles in the beer possess an overall negative charge. This is particularly true of yeast cells that possess a very significant negative charge as long as they are still alive. Hence, the collagen molecules act as a coagulant, neutralising the beer particulates’ surface charge and allowing them to overcome repulsive forces. The collagen molecules are long macromolecules with a triple-helix structure that are able to bridge between the beer particulates. The consequence of these interactions is that the beer particulates form large flocs that are able to sediment quickly. Isinglass is employed usefully in both beer that is matured in the brewery and in beer that is matured in cask. It may be added to beer in the cold storage vessel. This has the effect of speeding up the process, reducing costs and increasing brewery capacity. There is also a superior partition of the solid and liquid phases. The bulk of the beer (the ‘supernatant’) will be faster and less expensive to process downstream, notably in the filtration operations. Also, flocculents are capable of removing very small particles that even filtration processes may allow to pass. Finings therefore have a beneficial effect on colloidal stability. As described earlier, isinglass finings are essential for the clarification of unfiltered cask-conditioned beer. The cask undergoes a suitable period of secondary fermentation. After placement in the retail outlet, the finings will have clarified the beer after typically settlement for 24 h. The finings suspension is viscous compared with the beer and also comprises only a very small fraction of the volume of the beer. Thus, there is a challenge to achieving a homogeneous mix to optimise performance. The finings are often added in-line during tank transfer, perhaps on a pipe bend to maximise turbulence or the turbulence of fill is used in cask. However, one study has identified an optimised mixing regime comprising a short, relatively intense mix followed by a longer, more gentle mix, ideally achieved by two static mixers (Freeman et al., 2003). The performance of isinglass finings may be enhanced by materials called auxiliary finings. These are polysilicates, polysaccharides or a mixture of both. Addition to beer is before the addition of isinglass, and best performance requires good mixing of the beer and auxiliary finings. Auxiliary finings form a negative charge at beer pH. Hence, they support the performance of © Woodhead Publishing Limited, 2010
Separation technologies in brewing 437 the isinglass by ensuring that the beer particulates adopt a negative charge. This is especially true for yeast cells; if the cells are dead, and typically a moderate percentage are, they lose their surface charge. An alternative to isinglass has been identified based on pectin (Duan et al., 2008). It is claimed to perform as well as isinglass, and would have the commercial advantage that the resultant beer would remain indisputably suitable for consumption by vegetarians.
15.8 Filter aid filtration and applications in brewing Filtration in breweries is most commonly accomplished by the use of filter aids. These substances, used as slurried powders, form incompressible and highly porous filter beds, thus allowing the relatively free flow of beer. In small plants with small batch sizes, simpler filtration technologies may suffice. The most common filter aid used in breweries is kieselguhr or diatomaceous earth (Fig. 15.2). These materials comprise fossils or skeletons of microscopic saltwater or freshwater life known as diatoms. When they die they sink and form deposits that are mined, processed and size-classified to give kieselguhr of various grades. The disadvantages of kieselguhr are that it is a health hazard (by dust inhalation) in its dry form as delivered to the brewery and that it is in itself non-biodegradable, with a concentration of organic solids, and is thus expensive to dispose of in landfill sites. The configuration of a filter aid filtration system is shown in Fig. 15.3. Before processing occurs a precoat of filter aid is deposited onto the filtration
Fig. 15.2 An electron micrograph of kieselguhr (diatomaceous earth).
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438 Separation, extraction and concentration processes Buffer tank
Filter aid slurry tank
Base of cold storage tank Dosing pump
‘Rough’ beer
Filter
Pump
Filtered beer
Recycle and precoat facility
Fig. 15.3 A diagrammatical representation of a filter aid filtration system.
surface. This is achieved by recycling of a water/filter aid slurry around the filter. After several minutes, the precoat is deposited completely onto the filtration surface and the recycling water is clean. The precoat is necessary to ensure efficient filtration of the early part of the beer run, to guarantee the integrity of the filter throughout the run and to aid removal of the filter cake after the process cycle. After precoating, the filter is smoothly put into ‘forward flow’ mode. The filter aid slurry is added continuously to the flowing beer stream. Thus, the filtration surface is constantly being regenerated. In this way, the filtration run time is extended causing the process to be commercially viable. With regard to the actual filter unit, options may be divided into plate and frame type, leaf type and candle type. Plate and frame filters have been a workhorse of the brewing industry since the inception of filter aid filtration. They are known to enable excellent filtrate clarity. However, they are not amenable to full automation causing long down time between filter runs and an increased manpower requirement. Most brewers would now choose leaf or candle filters for which beer recovery, cleaning and re-starting may be automated by process control systems. Selection of leaf or candle is a brewery specific decision. Leaf filters are mechanically more complex and higher maintenance, but have more flexibility in flowrate and are not as vulnerable to process interruption. The bulk filtration duty in a brewery is a demanding unit operation. It is essential for product clarity, and also for colloidal stability. It should significantly lower the quantity of contaminant micro-organisms presented to the pasteuriser, because heat should be used sparingly if flavour impairment is to be avoided. If sterile filtration is employed, the bulk filtration stage must still give a high degree of clarity because the majority of sterile filtration systems have very limited solids-holding capacities. In this chapter, alternative filtration technologies and filter aids are
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Separation technologies in brewing 439 considered. These alternatives are challenged by the intrinsic difficulties of beer as a product for filtration. The low temperature (0 °C) and presence of dissolved solids and alcohol means that viscosity is quite high (at least 2 mPa s). Of even more significance is the nature of the suspended solids. These may be present in very high levels, perhaps up to 0.2% by volume or even higher over short periods during tank run-off (Fig. 15.4). Practically all of the suspended beer solids are compressible, which causes them to form filter cakes impermeable to beer flow. Many are small, meaning that filtration must be through fine, small flow channel media to achieve acceptable clarity. Filtration may also be impaired by colloidal substances such as b-glucan gels (Waiblinger, 2002). Alternative filter aids (and alternative technologies) are currently in use. Perlites consist of thermally expanded volcanic glass (Davies, 2004), crushed to form microscopic flat particles (Fig. 15.5). Perlites are less efficient filter aids than are kieselguhrs but are perceived as being safer than kieselguhr. However, it lacks the remarkable skeletal structures of the diatoms that comprise kieselguhr. As a consequence its filtration performance is not as good. In order to achieve the required filtration performance, secondary filtration (e.g. sheet or cartridge filters) is required. It is interesting to compare the filtration performance of kieselguhr and perlite in some detail (Fig. 15.6). This graph shows the particle sizes remaining in beer after filtration with kieselguhr and perlite (of similar permeability). Note that the kieselguhr exhibits a very exact particle size at which almost all smaller particles will pass through and almost all larger particles will be retained. This is not the case with perlite, which although it removes more very small particles than Yeast cells
25 20 15 10 Protein-polyphenol particles 5 0 0.5 0.6 0.6 0.7 0.8 0.9 0.9 1.0 1.1 1.3 1.4 1.5 1.7 1.8 2.0 2.2 2.5 2.7 3.0 3.3 3.6 4.0 4.4 4.9 5.4 5.9 6.5 7.2 7.9 8.7 9.6 10.6 11.7 12.9 14.2 15.6 17.2
Volume in size class (mm3 ml–1) ¥ 10–6
30
Particle diameter (microns)
Fig. 15.4 The particle-size and concentration distribution in a sample of pre-filter beer.
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440 Separation, extraction and concentration processes
Fig. 15.5 An electron micrograph of perlite.
Volume of particles in class (ml l–1)
4.5 4.0 3.5 3.0 2.5 2.0 1.5 1.0 0.5
16.0
10.1
(mm)
6.4
Particle d iameter
4.0
2.5
1.6
1.0
0.6
0.4
0.0
Unfiltered beer Perlite filtered Kieselguhr filtered
Fig. 15.6 Particle-size distributions and concentrations of samples of beer before filtration and after filtration with perlite and kieselguhr.
kieselguhr, it allows particles in excess of 1.5 mm to pass into filtered beer. The two filter aids operate in a very different manner. The kieselguhr behaves somewhat like a sieve. The perlite, however, filters more like a depth filter (mass filter). The performance of perlite shown in Fig. 15.6 is less desirable than the kieselguhr filtration. The larger particles will manifest themselves as a more obvious, visible haze in the beer. Also, a higher proportion of any
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Separation technologies in brewing 441 spoilage bacteria will pass through the perlite filter. Hence, there is a need for secondary filtration or an increase in pasteurisation intensity. The latter may impinge adversely on beer quality. It is likely that many of the novel, alternative filter aids now on the market will not filter beer to the same standard of clarity that is possible with kieselguhr. This is because of the unique structures in kieselguhr that cannot economically be reproduced artificially. The advantages of the alternative filter aids must be one or more of: a stabilisation effect, the ability to be regenerated and reused, and improved health and safety or environmental factors. After the main filtration operation, many breweries employ secondary filtration operations. Often the purpose is to provide sparkling clarity and enhanced shelf life before the beer becomes turbid in-pack. Some products are sterile filtered. This means that essentially all microbes with the potential to spoil beer are removed. The most common technologies installed for these purposes today are sheet filters (Brunner, 1987) and cartridge filters (Tubbs, 1998). Sheet filters mainly consist of cellulosic fibres that have been compressed into a thin mat and arranged in a plate and frame filter press. Cartridge filters comprise small units enclosing a filter element commonly of polymeric fibre sheets that are pleated. The demands on beer quality today, especially if sterile filtration is required, mean that, on occasions, three or four filtration steps are performed.
15.9 Regenerable and reusable filter aids and applications in brewing It is clear that kieselguhr would become much more environmentally acceptable if it were to be reused in the process. Alternatively, the spent filter cake could become a useful co-product of the brewing process. It has been suggested that spent kieselguhr could be employed for enhancing the nutrient value and structure of agricultural soil either directly or by composting first (Russ, 1993). However, in recent years, this use has become unfashionable because of the perceived health risk. It is possible to add spent kieselguhr to construction materials such as bricks and tiles. The problem here is economic in that individual breweries do not produce enough to justify the transportation to central processing. Hence, there is still scope for engineers to develop processes that enable kieselguhr to be reused in the brewery. There are established technologies involving sodium hydroxide and also furnacing (Russ, 1993). In both cases, economy of scale suggests the need for removal of spent kieselguhr from the brewery to a centralised processing station. Recently, studies have been performed on the use of hydrocyclones to separate the filter aid particles from organic material (Poku, 2004). Hydrocyclones operate by the conversion
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442 Separation, extraction and concentration processes of pressure energy into vortex flow, thus enabling separation of relatively dense particles from liquid suspension and less dense particles. Filter aid particles may be concentrated by a factor of up to twenty-five (depending on particle size) in one pass. Yeast cells are not concentrated and therefore are effectively washed out from the spent filter aid. There is the possibility of utilising several in series to effect a good separation, but there will still be a need for chemical washing of some stubborn organic residues. With all the available technologies the skeletal structure of kieselguhr means that only partial recycle is possible. There will be loss of material owing to particle attrition. There have been a variety of synthetic, regenerable filter aid filtration systems proposed. One system is based on a synthetic polymer (Brocheton et al., 1995). It is granular with a typical particle size of 35 mm, which is larger than the kieselguhr grades normally employed. However, the particles are claimed to be hydrodynamically ‘lighter’ than kieselguhr, which assists in the development of smooth, even filtration cakes. Regeneration is by hot caustic solutions, this means that the filter aid may be blended with polyvinyl polypyrrolidone (PVPP) and regenerated together. PVPP is a polyamide that adsorbs the polyphenolic component in beer that is responsible for combining with certain polypeptide fractions and polymerising into visible particles. Thus, it extends beer shelf-life. Stabilisation (of polyphenolic sources of instability) and filtration are achieved in one unit operation. Another system employs a mixture of synthetic microballs (for filtrate clarity) and fibres (for cake flexibility) (Harmegnies et al., 1997). Similarly to the previous system, the main regeneration process is with a hot caustic solution and therefore PVPP may be incorporated into the filter aid to provide stabilisation and filtration. A major chemical company has recently commercialised a regenerable filter aid comprising 30% by mass PVPP on a matrix of polystyrene (Zimmermann et al., 2008). In order to maintain the permeability of the filter bed, the particle size of the particles is larger than that of an equivalent kieselguhr. Conceptually, this would result in poorer filtered beer clarity, but the fact that the material is regenerable justifies a deep precoat layer and higher bodyfeed dosage rate enables acceptable beer clarity. Employment of candle filtration technology has been recommended with this filter aid (Ferstl and Zuber, 2009). With all three systems, filtrate quality and an evaluation of economic worth were both positive. Kieselguhr suffers from the disadvantage of transmission of oxidising transition metal ions into beer (notably iron) that the regenerable filter aids do not. Disadvantages of regenerable filter aid systems are the need for precoating and bodyfeed mixtures to be identical and also there is very limited flexibility over the dosage rate of stabiliser if it is employed. The economic benefits would depend on the current process. The required process modifications may be minor or substantial, thus affecting the economic attractiveness of the process. Pilot trials are recommended to assess process variables such as dosage rate (and consequently run length) and frequency of regeneration.
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Separation technologies in brewing 443
15.10 Bulk beer filtration by membranes In recent years, an acceptable alternative to filter-aid filtration has been developed that should reduce costs and improve the environmental image of brewers. This technology is cross-flow (or tangential flow) microfiltration by membranes. The challenge facing membrane filtration is the presence of particles in the micrometre and sub-micrometre ranges (Fig. 15.4) that cause severe membrane fouling. A partial solution is to pump the unfiltered beer (known as retentate in cross-flow processes) across the membrane surface, inducing a scouring effect. However, in practice, this is only partially successful because of the presence of small particles, colloids and macromolecules (Fig. 15.7). Nevertheless, the technology has been developed to the extent that the process economics are now viable. Currently, there are three commercially available systems for primary filtration of large volumes of beer. Norit supply a tubular membrane system (Noordman et al., 1999). The cost of replacement membranes is an important operating cost, and the Norit membranes have been shown to last for over 500 filtrations. The flow rate performance of the membranes is improved by several ‘flux-enhancement’ techniques. As the flowrate through a module falls, the pressure drop across the membrane is gradually increased, thus increasing the fouling rate and shortening the filter run, but maintaining an acceptable flow rate. Another technique is to employ ‘back flushing’. This entails increasing the pressure on the filtrate (known as permeate in crossflow processing) so that it exceeds the pressure of the retentate for a few seconds every, say, two hours. Although there is therefore a loss of acquired permeate, the effect of the backflush is to remove some of the fouling layer back into the retentate. This temporarily improves the flow rate of permeate after backflush. Norit have also developed cleaning regimes that include enzymes as well as detergents.
Molecular gel layer Yeast
In depth pore plugging Macromolecules
Cake formation
Surface adsorption Fine particles
Fig. 15.7 Diagrammatical representation of the membrane fouling mechanisms in crossflow microfiltration of beer.
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444 Separation, extraction and concentration processes The ‘Profi’ system is supplied by a collaboration between Westfalia (part of GEA) and Pall Seitz Schenk (Anon., 2004). Again, the membrane format is tubular. However, this system is characterised by a high-performance disc-stack Westfalia centrifuge immediately upstream of the membranes. The centrifuge increases the membrane flow rate, but, because the fouling is mostly caused by small particles that are inefficiently removed by centrifugation, the performance benefit is actually quite small (0–50%). The main benefit is reduction of the final volume of retentate before the suspended solids concentration is too high to allow efficient processing. Essentially, beer recovery from slurry operations has been moved upstream of filtration. Indeed, it is claimed that, in at least some installations, there is little or no recycling of retentate (Anon., 2009). The membranes are regenerable with caustic and oxidising detergents, and are sanitisable up to 80 °C. A third system is supplied by another collaboration between Alfa Laval and Sartorius (Borremans and Modrok, 2003). Again, the membranes are preceded by a high-performance (Alfa Laval) centrifuge. However, the Sartorius membranes are in a very different format to the systems described above. They are supplied as cartridges of square membrane flat sheets. They are separated on the retentate side by turbulence promoting gauze, which helps the cross-flow effect but necessitates the use of a centrifuge to prevent blocking of the cartridges on the retentate side. Cleaning is claimed to only require rising with water and caustic solution. The membranes comprise a porous structure that is asymmetric. For cross-flow processes, this is optimal when the narrow pore diameter is on the retentate side. The benefit of using the centrifuge is debatable. With the latter two systems, the power supplied to operate the centrifuge is at least partly recovered by a lesser need for a high cross-flow velocity over the membranes and a consequent reduced refrigeration demand. There will be a small increase in temperature by centrifuge processing, perhaps 1 °C, which is potentially slightly damaging to the colloidal stability of the beer owing to re-solution of some of the beer haze particles. It is worthy of note, however, that centrifuges have improved significantly in recent years, with improved particle removal performance and increasingly sophisticated and reliable sealing systems (Meckler, 2003). One of the main developments that has enabled improved performance is the employment of polyether sulfone (PES) membranes. These were until recently difficult to manufacture with the required microporous structure and with sufficient repeatability, but owing to an important development this is no longer the case (Riddell, 2002). For the filtration of beer they exhibit excellent low-fouling properties. The disadvantage of polymeric membranes compared with ceramic membranes, for example, is that replacement is required more frequently. The polymeric membranes last probably at least 500 hundred process runs, perhaps equivalent to two years operation. All three of the main commercially available systems employ PES membranes. The cross-flow aspect of the technology makes the plant more viable on a
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Separation technologies in brewing 445 large scale. On a small scale, membrane filtration is commonly employed for sterile or polishing filtration, often with a cartridge filter. The membranes are supplied in modular form. This has the benefit that membrane cartridges may easily be replaced when required without interruption to process. Also, some modules may be undergoing cleaning while others are processing beer. This raises the possibility of continuous processing. Cross-flow microfiltration is more amenable to automation than filter aid filtration and so there is the potential for a continuous process that requires minimal process supervision. However, a recent installation has shown more need for supervision than was anticipated (Pickerell and Heeb, 2008). Operator experience has been required in the operation of the centrifuge, control system and instrumentation maintenance, although this is an area for improvement and development in the future. Beer stability is very sensitive to the presence of oxygen as a source of free radicals that cause the beer to taste stale (oxidised) and have a deleterious effect on colloidal stability and thus shelf-life. A major cause of oxidative damage are the ions of transition metals, notably iron and manganese, with iron often employed as a marker for undesirable concentrations. A disadvantage of the most commonly employed filter aid, kieselguhr or diatomaceous earth, is that impurities in the silica structure solubilise in beer. Although the increase in dissolved iron is small, this can easily be the source of half of the concentration in the beer. The elimination of this effect by the absence of kieselguhr may be of significant benefit to product quality. Reduction in free radical formation through the employment of membranes instead of kieselguhr has been measured (Broens et al., 2007). Processing costs are generally less sensitive to the filterability of the beer for membrane filtration than filter aid filtration. The relative compactness of the plant means that cleaning is more rapid and there is less loss of product. Water usage is reduced, although the need to regenerate fouled membranes means that there is increased usage of detergents. There is scope to employ the membranes as sterile filters, negating the need for pasteurisation or secondary sterile filtration. However, current systems are mostly employed only as primary filters. This is because many brewers employ stabilisation processes downstream of filtration. Also, sterile filtration would mean that each module would require integrity testing (confirmation that the membrane pores remain small enough for sterilisation) before each filter run. There is also the possibility that membranes could fail midrun, although there is instrumentation that employs advanced light-scatter techniques that are claimed to detect membrane failure (Wilhelm, 2009). There is likely to be enormous potential for development of improved membrane process efficiency. For example, membrane manufacturing techniques derived from the electronics industry have been employed to manufacture so-called microsieves (Lommi et al., 2003). These comprise silicon nitride wafers. They feature a very high porosity, very precise and consistent pore diameter and membrane thickness of the order of just 1 mm.
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446 Separation, extraction and concentration processes This enables a requirement for very low trans-membrane pressures, and also the application of a high frequency backflush (‘backpulse’) to maintain permeate flow rate. Consequently, the membranes achieve a permeate flow rate per unit area with beer of the order of a hundred times greater than with PES membranes. However, the current absence of mass production that would allow economical full-scale plants, and also problems optimising the backpulse in larger plants, have as yet prevented successful commercialisation of the technology.
15.11 Recovery of cleaning detergents in brewing Operation of modern breweries is characterised by the application of cleaningin-place (CIP) technology. This is characterised by a cycle of recirculating water rinses, detergents and disinfectants in the appropriate sequence. Advantages include effective automation, reliable programmable cleaning cycles and the ability to maintain a closed process. Water economy is achieved by reusing final rinses as first rinses and automation of the disposal of spent cleaning agents and automated top-up of cleaning agent strength. Nevertheless, a single typical cleaning cycle employs a further 10% of the water that finishes in the product (Freeman, 2008). However, application of nanofiltration membrane technology enables removal of fouling from the detergent. The detergent, either acidic or alkaline, is recovered at the rate of typically 90%. It is mostly the relatively small detergent molecules that pass through the membrane, but there is some transmission of soil so that some system purging is required. The economic viability of the installation is improved with economy of scale. As with other membrane processes, viability is likely to improve in the future as membrane performance is improved and membrane prices fall in real terms. Also the prices of water, water treatment and detergents are certain to increase in the future. Currently, dependent on process conditions and geographical location, payback time for capital expenditure has been calculated as less than two years (Catala and Freeman, 2008). A more recent calculation suggests that payback time will actually be typically less than one year (Freeman, 2008). Future improvement to the technology could be automation based on in-line measurement of contamination in the detergent, perhaps based on absorption of near infra-red radiation.
15.12 Dissolved gas control by membrane technology As a carbonated beverage, the concentration of carbon dioxide in beer is one of the most important factors contributing to its flavour. In a modern brewery, most typically, carbon dioxide has to be added to attain the relatively high
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Separation technologies in brewing 447 specification of most small-pack or keg beers. This is often achieved by inline post-filtration, with a feedback control loop and controlled dosage of carbon dioxide into a turbulent zone that ensures dissolution. However, on occasions there is a need to reduce carbon dioxide concentration to bring it within the specification. This is problematical, because it usually requires a nitrogen purge that risks loss of foam potential in the product and can form unsightly particles as the foam collapses. Furthermore, the need to reduce carbon dioxide concentration may also occur because the beer is a mixed gas or ‘nitro-keg’ product. These products exhibit a ‘smooth’ palate, enhanced foam potential with an attractive ‘theatre in dispense’ (Cooper, 2009). Thus there is scope for a process that can adjust gas concentrations in beer without dispersing bubbles. A suitable technology relies on hydrophobic membranes (Gill and Menneer, 1997). These comprise hollow membrane fibres made of a hydrophobic polymer, commonly polypropylene. The membrane pores do not allow passage of the beer up to a certain pressure. To achieve the pressure differential required in brewery processing, the pores have to be as small as 50 nm in diameter. The beer is circulated on one side of the fibre (most probably the tube side) and a particular mixture and pressure of gases is applied to the other side. A vacuum will allow a reduction of carbonation. Application of nitrogen allows nitrogenation of the beer whilst simultaneously reducing carbonation, because each gas diffuses according to its own partial pressure. A further advantage is that because there is no oxygen in the applied gases, there is potential to reduce the residual oxygen content of the beer, assisting product stability. The technology has been successfully applied in UK breweries making ‘nitro-keg’ beers (Cooper, 2009).
15.13 Future trends As with all industries, future business strategies are required to pay greater attention to environmental issues such as water and energy economy and waste reduction. There is the prospect of standardisation of calculation of water and energy footprints that will become increasingly important key performance indicators. It appears that the brewing industry is partitioning into two distinctive types of company: major international brewing companies that market ‘key brands’ globally and local concerns that produce perhaps more ‘niche’ products. In the former instance especially, there is scope for employing key performance indicators as marketing tools to gain advantage over brand rivals. It is possible that there is a significant move towards beer products of relatively low alcoholic strength (Evans, 2008). This would enable beer to be produced by still brewing to high alcoholic strength (‘high gravity brewing’) but increasing rates of dilution near the end of the process, thus economising on utility usage. © Woodhead Publishing Limited, 2010
448 Separation, extraction and concentration processes Rising utility and effluent treatment costs may result in more widespread introduction of sophisticated processing of effluent streams. For example, anaerobic digestion may be employed to generate biogas fuel from effluent streams. The residual effluent could be further sequentially processed culminating in reverse osmosis. The resultant water is effectively completely clean, nearly pure, and could be reused as a brewing raw material.
15.14 References Anon. (2004), Official inauguration of kieseguhr-free filter line at Pott’s Brewery in Oelde/Westphalia (Germany), Brauwelt International, 22(5), 372–373. Anon. (2009), Greenfield brewery with kieselguhr free filtration technology, Filtration, 9(3), 169. Bamforth C W (2000), Making sense of flavour change in beer, Technical Quarterly of the Master Brewers’ Association of the Americas, 37(2), 165–171. Borremans E and Modrok A (2003), Membrane filtration by Alfa Laval and Sartorius, Brewer and Distiller International, 34(4), 10–11. Boulton C (2009), Stirring stuff. Getting the best out of cylindroconical fermenters, Brewer and Distiller International, 5(6), 18–21. Broens L, Meijer D, Mepschen A, Schuurman R, Methner F, Kunz T, Eisenblatter F, Metz L and Brunacker J (2007), Practical membrane filtration for beer clarification, Proceedings of the European Brewery Convention, Venice. Brocheton S, Hermia J, Rahier G and van den Eynde E (1995), The basic principles of a new beer filtration process, Proceedings of the European Brewery Convention Brussels, Nurnberg, Fachverlag Hans Carl, 427–436. Browne J (2008), Innovation is about survival. The challenges of a modern small pack line, Brewer and Distiller, 4(1), 11–17. Brunner M (1987), Filter sheets, International Bottler and Packer, 61(6), 27–32. Catala M and Freeman G (2008), The relationship between water consumption and energy usage in the malting and brewing industries: opportunities and priorities, Proceedings of the World Brewing Congress 2008. Cooper D (2009), The nitro-keg revolution, Beers of the World, 24, 54–56. Davies K (2004), Filter aids, The Brewer International, 4(2), 14–18. Duan D, Rogers P, Dawson J, Aspridis C, Day G, Delaere S and Oliver T (2008), The use of pectin-based finings in commercial-scale beer making, Proceedings of the 30th convention of the Asia Pacific Section of the Institute of Brewing and Distilling, Auckland. Evans E and Kaur M (2009), Keeping sleepy yeast awake until bedtime, Brewer and Distiller International, 5(5), 38–40. Evans J (2008), Low alcohol beers reach new heights, Brewers’ Guardian, 137(10), 29–32. Eyben D, Hermia J, Meurens J, Rahier G and Tigel R (1989), Industrial results of a new wort filter, Proceedings of the 22nd Congress of the European Brewery Convention, Zurich, 275–281. Ferstl FF and Zuber J (2009), Pre-coat filtration with a new, regenerable filter aid, Proceedings of the 13th Scientific and Technical Convention of the Institute of Brewing and Distilling Africa Section. Freeman G (2008), Cleaning-in-place, Campden BRI’s ‘Cleaning and disinfection conference – managing new challenges’. Freeman G J, Powell-Evans M H B, Baron J M, Dawson M K, Patel A, Skipper A J, Evans C T, Boulton C A, Grimmett C M and Le Gourrierec X (2003), Improving the © Woodhead Publishing Limited, 2010
Separation technologies in brewing 449 effectiveness of isinglass finings for beer clarification by optimisation of the mixing process part 3: full-size prototype evaluation, Journal of the Institute of Brewing, 109(4), 326–331. Gill C B and Menneer I D (1997), Advances in gas control technology in the brewery, The Brewer, 83(987), 77–84. Harmegnies F, Bonacchelli B and Tigel R (1997), Beer filtration with regenerable filter aid, Proceedings of the European Brewery Convention Maastricht, Nurnberg, Fachverlag Hans Carl, 517–524. Harmegnies F, Bonacchelli B and Tigel R (2003), Continuous brewing, Proceedings of the 29th Congress of the European Brewery Convention, Dublin. Lake J C and Speers R A (2008), A discussion of malt-induced premature yeast flocculation, Technical Quarterly of the Master Brewers Association of the Americas, 45(3), 253–262. Lommi H, Raspe O J, van Rijn C J M and Vos J (2003), New filter and method for beer clarification, Proceedings of the 29th European Brewery Convention Congress. Malfliet S, Van Opstaele F, De Clippeleer J, Syryn E, Goris K, De Cooman L and Aerts G (2008), Flavour instability of pale lager beers: determination of analytical markers in relation to sensory ageing, Journal of the Institute of Brewing, 114(2), 180–192. Martin I (2002), Pasteurisation possibilities for plastics, Brewers’ Guardian, 131(1), 20–22. Meckler O (2003), New generation centrifuges in breweries, Proceedings of the 9th Brewing Convention of the Institute and Guild of Brewing Africa Section, Victoria Falls, Zambia, 178. Noordman T R, Berghuis O A E, Mol M N M, Peet C J, Muller J L M, Broens L and van Hoof S (1999), Membrane filtration for bright beer, an alternative to kieselguhr filtration, Proceedings of the 27th Congress of the European Brewery Convention, 815–822. O’Rourke T (1999), Mash separation: a review, Brewers’ Guardian, 128(2), 15–16. O’Rourke T (2000), Microbiological quality, Brewers’ Guardian, 129(10), 39–42. Pickerell A and Heeb W (2008), Cross-flow membrane filtration at the Coors Shenandoah brewery, World Brewing Congress 2008 Proceedings. Poku M (2004), An investigation into the recycling of filter aids for the brewing industry, PhD Thesis, University of Essex Department of Biological Sciences. Riddell P (2002), Sterile filtration of beer. The PES story, Brewer International, 2(11), 31–35. Russ W (1993), Disposal of kieselguhr – kieselguhr recycling, Brauwelt International, 11(1), 51–55. Siebert K J (1997), Beer clarity stability, Proceedings of the 6th Central and Southern Africa Section Convention of the Institute of Brewing, 67–78. Soares E V (2009), Flocculation in Saccharomyces cerevisiae, in Preedy V R, Beer in health and disease prevention, London, Elsevier, 103–112. Tubbs J (1998), Cartridge filtration – part 2, The New Brewer, 15(3), 56–58. Virden J (1995), How to avoid ‘mincing’ your trub, BDI, 26(11), 20–21. Waiblinger R (2002), Beer filterability, The Brewer International, 2(1), 15–18. Wilhelm M (2009), Continuous integrity monitoring of membrane filter units, Brauwelt International, 27(3), 143–145. Zimmerman T and Meffert M (2008), Crosspure® – The future of kieselguhr-free beer filtration, Proceedings of the 30th Convention of the Asia Pacific Section of the Institute of Brewing and Distilling, 1 p.
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450 Separation, extraction and concentration processes
16 Methods for purification of dairy nutraceuticals C. J. Fee, J. M. Billakanti and S. M. Saufi, University of Canterbury, New Zealand
Abstract: The properties of the classes of proteins found in bovine milk and increasingly used as nutraceuticals are reviewed. Whey proteins, an important class of dairy nutraceutical products, can be classified as acidic and basic proteins and immunoglobulins. The methods used for their purification, including ion exchange, chromatography and membranes are described. Key words: nutraceuticals, milk proteins, lactoferrin, ion exchange, chromatography, membranes.
16.1 Introduction The global market for nutraceuticals was worth US$117bn in 2007, US$124bn in 2008 and is expected to reach US$177bn by 2013 at a compound growth rate of 7.4% per year, driven in part by increased demand from developing countries. Many biologically active components and derivatives of milk and their nutraceutical applications have been reviewed by Severin and Xia (2005). These include minor milk proteins such as lactoferrin, lactoperoxidase, immunoglobulins and lysozyme, endogenous peptides and those derived from protein hydrolysis, oligosaccharides, hormones, growth factors and gangliosides. Of the proteins, it is mainly the whey proteins that have nutraceutical applications. However, casein micelles have been proposed as a novel delivery vehicle for nutraceutical compounds, taking advantage of the casein protein self-assembly (Semo et al. 2006). In addition, milk proteins, including caseins, are a source of biologically active peptides that are inactive within the sequence of a native protein but can be released by enzymatic hydrolysis.
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Methods for purification of dairy nutraceuticals 451 Such bioactive peptides have been found to exhibit various physiological activities such as antihypertensive, opioid, immunomodulatory antimicrobial, antioxidative, antithrombotic, and cytomodulatory activities and may find use in the treatment of diarrhea, hypertension, thrombosis, dental caries, oxidative stress, mineral malabsorption, and immunodeficiency (Haque et al. 2009). Lactose, the main carbohydrate in milk, can be extracted and crystallized, and its derivatives used as nutraceuticals (Ganzle et al. 2006). However, because enzymatic hydrolysis is generally required to produce bioactive peptides and chemical modification is required to produce bioactive lactose derivatives, rather than extraction/purification of either class of nutraceuticals in their native form, we will not consider these products further here. The milk fat globule membrane has recently been shown to contain a rich variety of components that have an impressive array of functionalities, including cholesterolemia-lowering factor, inhibitors of cancer cell growth, vitamin binders, inhibitor of Helicobacter pylori, inhibitor of b-glucuronidase of the intestinal Escherichia coli, xanthine oxidase as a bactericidal agent, butyrophilin as a possible suppressor of multiple sclerosis, and phospholipids as agents against colon cancer, gastrointestinal pathogens, Alzheimer’s disease, depression, and stress (Spitsberg 2005). However, these components are present at extremely low levels, so their economic fractionation would be challenging. The basic whey proteins, lactoferrin, lactoperoxidase and lysozyme, have known antimicrobial properties, whereas the main acidic whey proteins have functions that include lactose synthesis, Ca2+ transport, immunomodulation and anticarcinogenicity (a-lactalbumin) and retinol transport, fatty acid binding and antioxidant activity (b-lactoglobulin). Lactoferrin, in particular, has multiple activities, including antioxidant activity, iron binding, anticarcinogenicity and immunomodulation (Severin and Xia 2005). Immunoglobulins, unsurprisingly, impart various immunoprotective functions in a range of functional foods (Gapper et al. 2007). Table 16.1 shows the main components of milk, their concentrations and their bioactivities relevant to nutraceuticals (Severin and Xia 2005). In this chapter, we focus on the purification of whey proteins as nutraceuticals. The main workhorses of whey protein purification are ionexchange chromatography and membrane separations. However, examples of affinity chromatography, adsorptive membranes and other separation formats aimed at large-scale processing are also presented. We have organized the sections by product (acidic proteins, basic proteins and immunoglobulins) and describe examples for the separation/purification of each.
16.2 Components of acidic whey protein 16.2.1 b-Lactoglobulin: properties and applications b-Lactoglobulin (b-Lac) is the major protein component in bovine whey, constituting approximately 58% of the whey protein or 10% of total milk © Woodhead Publishing Limited, 2010
452 Separation, extraction and concentration processes Table 16.1 Milk proteins and their bioactivities (adapted from Severin and Xia, 2005) Concentration (g L–1) Protein Total caseins
Human
Cow
2.7
26.0
a-Casein
13.0
b-Casein
9.3
k-Casein
3.3
Total whey protein
6.3
Ion carrier, bioactive peptides precursors
6.3 3.2
b-Lactoglobulin
Bioactivity
a-Lactalbumin
1.9
1.2
Immunoglobulins (A, M, G) Serum albumin
1.3
0.7
0.4
0.4
Lactoferrin
1.5
0.1
Lactoperoxidase Lysozyme
0.1
Miscellaneous
1.1
Retinol carrier, fatty acid binding, possible antioxidant Lactose synthesis, Ca2+ carrier, immunomodulation, anticarcinogenic Immune protection
Antimicrobial, antioxidative, immunomodulation, iron adsorption, anticarcinogenic 0.03 Antimicrobial 0.0004 Antimicrobial, synergies with immunoglobulins and lactoferrin 0.8
Proteose-peptone
1.2
Glycomacropeptide
1.2
Antiviral, bifidogenic
protein (Lozano et al. 2008). The concentration of b-Lac in whey is in the range 2–4 g L–1 (Andersson and Mattiasson 2006). The primary structure of b-Lac consists of 162 amino acids and it has a molecular weight (MW) of approximately 18.4 kDa. Six genetic variants are known, the most common being the A and B variants. The A and B variants differ in only two amino acid residues, at positions 64 and 118, which are Asp and Val for b-Lac A and Gly and Ala for b-Lac B (Elofsson et al. 1997). Because b-Lac A has an additional negative charge, it has slightly lower pI value (pI 5.1) than b-Lac B (pI 5.2), although the molecular weights of the variants are essentially the same (Yamamoto and Ishihara 1999). The secondary structure of b-Lac comprises nine strands of b structure, a short a helix segment and three helical turns. Its quaternary structure depends on the medium pH: it occurs mainly as a stable dimer, with a molecular weight of 36.7 kDa, at pH values between 5.2 and 7; as an octamer, with © Woodhead Publishing Limited, 2010
Methods for purification of dairy nutraceuticals 453 a molecular weight of 140 kDa, at pH values between 3.5 and 5.2; and as a monomer, with two-cysteine residues per monomer, at pH below 3.0 and above 8.0 (de Wit 1989). b-Lac is a good source of essential amino acids and has a potential use in power drinks owing to its good solubility (Horton 1995). Good gelling formation and superior foam stability compared with other whey proteins make b-Lac suitable for confectionery production (Cowan and Ritchie 2007; Zydney 1998). Madureira et al. (2007) summarized a number of biological functions of b-Lac: it plays a role in regulation of mammary gland phosphorus metabolism, as a transporter for vitamin D, cholesterol and retinol, the transfer of passive immunity to the newborn and enhancement of pregastric esterase activity (Madureira et al. 2007). The b-lac content of bovine milk is much higher than that of human milk (El-Agamy 2007; Fox and McSweeney 1998) and this has been identified as a potential source of allergic reactions to infant formulae seen in some children (El-Agamy 2007; Monaci et al. 2006; Suutari et al. 2006). Therefore, the removal of b-Lac from whey may broaden the applications of whey product derivatives in the food industry (Casal et al. 2006). b-Lac-free whey may serve as the primary protein constituent of hypoallergenic infant formulae that have protein compositions that are more similar to that of human milk (Casal et al. 2006). 16.2.2 a-Lactalbumin: properties and applications a-Lactalbumin (a-Lac) is the second largest protein component of bovine whey, comprising approximately 3.4% of total milk protein or 20% of whey proteins. The concentration of a-Lac in whey protein is 1.2–1.5 g L–1 (Andersson and Mattiasson 2006). On the other hand, a-Lac is the predominant whey protein in human milk, with a concentration of 2.44 ± 0.64 g L–1 (after day 30 of lactation) (Jackson et al. 2004). a-Lac is a strong Ca2+-binding protein, consisting of 123 amino acids in a single peptide chain with four disulfide bonds. Its molecular weight is about 14.2 kDa and it has a pI value of 4.2. Human and bovine a-Lac have approximately 76% fully conserved residues (93 out of 123 amino acids) (Chatterton et al. 2006). a-Lac is the preferred protein source for infant formulae, owing to its high tryptophan content, high digestibility and lower potential for causing allergies compared with b-Lac (Gurgel et al. 2000; Zydney 1998). Additionally, because of its high tryptophan content it is applicable as a nutraceutical and because of its high cytotoxicity it possesses therapeutic uses (Konrad and Kleinschmidt 2008). a-Lac also has strong affinity for glycosylated receptors on the surface of oocytes and spermatozoids and may thus have potential as a contraceptive agent (Zydney 1998). Its reported biological functions include cancer prevention, lactose synthesis and treatment of chronic stress-induced diseases (Madureira et al. 2007).
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454 Separation, extraction and concentration processes 16.2.3 Bovine serum albumin: properties and applications Bovine serum albumin (BSA) exists in whey at concentrations between 0.3 and 0.6 g L–1, its molecular weight is about 69 kDa and it has a pI value of around 4.7–4.9 (Andersson and Mattiasson 2006). It consists of 582 amino acid residues, with 17 intramolecular disulfide bonds and a single free thiol at residue 34 (Burr 2001). Biological functions of BSA include fatty-acid binding, an antimutagenic function and cancer prevention (Madureira et al. 2007). BSA also has good gelling properties (Matsudomi et al. 1991) and is widely used in food and therapeutic applications (Zydney 1998).
16.3 Purification technologies for acidic whey proteins 16.3.1 b-Lactoglobulin purification Extraction of b-Lac from whey has been achieved by various technologies, including chromatography, membrane filtration, selective chemical and thermal precipitation, membrane chromatography, foam fractionation and colloidal gas aphrons. A number of recent examples of acidic protein purification are described briefly here, with many further examples from the past 10–15 years tabulated in Tables 16.2 to 16.5. Selective precipitation As described below, b-Lac has been selectively isolated from whey by forming a complex with chitosan (Casal et al. 2006; Montilla et al. 2007), addition of ammonium sulfate (Lozano et al. 2008), precipitation of a-Lac with sodium citrate (Alomirah and Alli 2004) and by peptic hydrolysis followed by membrane filtration (Konrad et al. 2000). b-Lac interacted reversibly with chitosan by electrostatic interaction, forming a precipitate at pH 6.2. The b-Lac was recovered by dissolving the precipitate in 100 mM sodium acetate at pH 9 to give a recovery of 90%, with a purity of 95% (Montilla et al. 2007). The isolated b-Lac maintained its native structure and the use of non-toxic chitosan may be of interest in industrial applications. b-Lac was separated from other whey proteins by precipitation with 50% ammonium sulfate (Lozano et al. 2008). The precipitate was dissolved and then again separated using 70% ammonium sulfate, leaving a supernatant solution enriched in b-Lac. After dialysis, lyophilization and reconstitution in water, final purification is carried out by weak cation-exchange chromatography. The total yield and purity of b-Lac from 3.5 L whey were 14.32 and 95%, respectively. Meanwhile, Alomirah and Alli (2004) recovered b-Lac from supernatant solution after a-Lac precipitation with sodium citrate. After several further steps (i.e. washing, centrifuge and dialysis), b-Lac was recovered with a purity ranging from 83 to 90%. The yield of the b-Lac isolate from this process was reported to be in the range 47–69%. © Woodhead Publishing Limited, 2010
Methods for purification of dairy nutraceuticals 455 Table 16.2 Additional examples of membrane separations of acidic whey proteins Year
Author
Target Protein source protein
2009 Metsamuuronen a-Lac and Nystrom b-Lac
Configuration/material
Whey powder and fresh whey
Flat sheet (2 ¥ 10–3 m2): 30 kDa and 100 kDa regenerated cellulose, 50 kDa polyaramide, 20, 30 and 50 kDa PES, 100 kDa PSF Flat sheet: 30 kDa regenerated Single 2009 Bhushan and b-Lac, cellulose (Millipore), 25 mm Etzel b-Lac, single GMP glycomacropeptide diameter (GMP), binary b-Lac and GMP, whey 2008 Konrad and Flat sheet: 100 kDa PES (0.093 a-Lac Concentrated Kleinschmidt rennet whey m2). Spiral wound: 150 kDa PSF (5.5 m2) 2007 Cowan and Flat sheet: 100 kDa PES (47 mm a-Lac, Single a-Lac, Ritchie diameter) b-Lac single b-Lac 2007 Almecija et al. a-Lac, Whey Tubular ceramic membrane (Clover Inside Ceram, TAMI b-Lac, Industries, France), 300 kDa, area BSA, 0.045 m2 LF, IgG 2006 Bhattacharjee b-Lac, WPC (obtained Flat sheet: 10 and 30 kDa PES et al. a-Lac after pretreating (56 mm diameter) the whey) 2004 Cheang and b-Lac, WPI (spiked with Flat sheet: 30 and 100 kDa Zydney regenerated cellulose, 50 cm2 a-Lac, some BSA) BSA 2003 Muller et al. 50 kDa Ceram Membrane (TAMI a-Lac WPC (2003a) Industries, France), average pore diameter 12 nm, area 0.045 m2 1998 Lucas et al. Tubular: inorganic membrane a-Lac, WPC (Carbosep, Tech-Sep), 10, 15, 50 b-Lac and 150 kDa, ID 0.6 cm, OD 1 cm, length 60 cm, area 0.0113 m2 WPI, whey protein isolate; WPC, whey protein concentrate; UF, ultrafiltration; MF, microfiltration; PES, polyethersulfone; PSF, polysulfone.
Konrad et al. (2000) compared several techniques for isolation of b-Lac from whey on a large scale. They developed a fractionation method consisting of peptic treatment of 10 000 L of whey, followed by membrane filtration. Three other methods compared were trichloroacetic acid precipitation, a salting out procedure and selective thermal precipitation. The yields of native b-Lac achieved by these four methods were 67.3, 44.9, 46.7 and 49.6%
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© Woodhead Publishing Limited, 2010
Year
Author
Target protein
Protein source
Materials/configuration
Mode of interaction/ligand
2008
Brochier et al.
b-Lac
Microfiltered whey
Mixed mode–hexyl amine
2008
Etzel et al.
WPI
Whey
2008
Etzel et al.
WPI
Whey
2006
Liang et al.
Whey
2004
Schlatterer et al.
b-Lac, a-Lac, BSA, IgG b-Lac
HyperCel™ column (Pall BioSepra), crosslinked cellulose, 90 mm, column volume 2.5, 5 and 10 mL Mono S column (Amersham Pharmacia Biotech Inc.), 2.38 L, 10 cm diameter SP Sepharose Big Beads (Amersham Biosciences), 5.34 L column volume, 20 cm diameter, 17 cm height Sephadex G-200, 2.6 cm ¥ 70 cm
2004
Turhan and Etzel a-Lac, WPI
2004
Rojas et al.
a-Lac, b-Lac
2004
Doultani et al.
2003
Neyestani et al.
a-Lac, WPI, LP, LF b-Lac, a-Lac, BSA
2002 2002
Vyas et al. Lan et al.
b-Lac Total protein
Macro-Prep ceramic hydroxyapatite (BioRad), 80 mm, 12 mm × 88 mm Lactic acid SP Sepharose Big Beads (Amersham whey Biosciences), 80 mL column volume Protein fraction Sephadex G-25® HR-10/10 (Amersham Bioscience) from ATPS Whey SP Sepharose Big Beads (Amersham Bioscience), 80 mL column volume Whey (1) Sephadex G-50 (Amersham Biosciences), 65 cm length, 1.6 cm diameter, 130 mL column volume (2) DEAE column (Amersham), column volume 5 mL Whey Calcium biosilicate particles Whey powder Diaion HPA25 (Sigma)
Cation exchange–methyl sulfonate Cation exchange–sulfopropyl (SP) Gel filtration
Whey
Cation exchange–SP Size exclusion Cation exchange, SP Size exclusion and anion exchange
Affinity, ligand – all-trans-retinal Ion exchange
456 Separation, extraction and concentration processes
Table 16.3 Additional examples of chromatographic separations of acidic whey proteins
2001
Gurgel et al.
a-Lac
WPI
2000
Tellez and Cole
b-Lac, a-Lac, BSA, IgG
Whey
Polyhydroxylated methacrylate (TosoHaas Peptide AF Chelate 650) Biogel A 0.5m and 5m (Bio-Rad, Hercules, Electrochromatography CA, USA), column 1: 1.5 cm diameter ¥ 30 cm height, column 2: 2.5 cm diameter ¥ 60 cm height
ATPS, aqueous two phase system; DEAE, diethyl aminoethyl cellulose.
Methods for purification of dairy nutraceuticals 457
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458 Separation, extraction and concentration processes Table 16.4 Additional examples of adsorptive membrane separation of acidic whey proteins Year
Author
Target protein
Protein source Membrane configuration/ Ligands materials
Whey, single b-Lac Whey, single b-Lac, a-Lac and BSA, binary b-Lac and BSA 2006 Bhattacharjee b-Lac, a-Lac Permeate from et al. two stage UF 2009 Saufi and Fee 2008 Goodall et al.
1996 Splitt et al.
b-Lac, a-Lac, BSA b-Lac, a-Lac, BSA
b-Lac, a-Lac, BSA
Whey
1996 Freitag et al. b-Lac, Whey a-Lac, BSA, IgG
Flat sheet ethylene vinyl Q alcohol Sartobind MA (Sartorius) Q, DIEA
Vivapure Q Mini-H (Vivasciences); 240 ml volume Sartorius cellulose based membrane adsorber: MA Q15 (area 15 cm2), MA Q100 (area 100 cm2), MA D15 (area 15 cm2). Synthetic copolymerbased membrane adsorber (area 1300 cm2) Sartorius cellulose-based membrane adsorber: MA Q15 (area 15 cm2), MA MA S15 (area 15 cm2)
Q Q, DIEA
Q, SP
Q, quaternary ammonium; DIEA, diethylamine; SP, sulfopropyl.
for the peptic treatment method, acid precipitation, salting out and thermal precipitation, respectively. The purity of b-Lac achieved by all methods was more than 90%. Chromatographic techniques Neyestani et al. (2003) used a series of chromatography steps to obtain pure b-Lac from whey after precipitation with 50% ammonium sulfate. Both the precipitate and supernatant solution obtained were dialyzed and lyophilized for further separation by chromatographic methods. A lyophilized precipitate fraction was reconstituted in distilled water and run onto a gel filtration column (Sephadex G-50, 131 mL column volume, 65 cm length) to obtain a first peak containing a mixture of BSA and casein and a second peak of pure b-Lac. The yield of b-Lac was 166 mg based on 50 mL of starting milk. Meanwhile, the lyophilized supernatant was dissolved in water and loaded onto a diethylaminoethyl cellulose (DEAE) anion-exchange column. Stepwise elution was applied to the column, resulting in two separate peaks consisting of a mixture of BSA and a-Lac in the first peak and pure b-Lac
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Methods for purification of dairy nutraceuticals 459 Table 16.5 Additional examples of other separation techniques for acidic whey proteins Year
Author
Technique
Target Protein Protein source
2009
Shea et al.
Foam fractionation
2008
Lozano et al.
a-Lac, b-Lac Spray dried whey powder, WPI Whey b-Lac
2005 2005
Fuda et al. Tolkach et al.
2004
2001
Alomirah and Alli Baumeister et al. Muller et al. Selective precipitation (2003b) Rodrigues et al. Aqueous two-phase system
2000
Konrad et al.
Ammonium sulfate precipitation 2008 Monteiro et al. Aqueous two-phase system a-Lac, b-Lac WPI 2007 Lucena et al. Acid precipitation Sweet whey, a-Lac WPC, WPI 2006– Casal et al.; Selective precipitation with b-Lac Rennet whey 2007 Montilla et al. chitosan
2003 2003
Colloidal gas aphrons Selective thermal denaturation of b-lac Chelating agent precipitation Expanded bed adsorption
Peptic hydrolysis and membrane filtration
b-Lac a-Lac
Whey WPC
a-Lac, b-Lac Whey, WPC, WPI Orotic acid
Whey
a-Lac
WPC
a-Lac, b-Lac Pure a-Lac, b-Lac, WPC Whey b-Lac
WPC, whey protein concentrate; WPI, whey protein isolate.
in the second. The yield of b-Lac in the second peak was estimated to be about 35 mg mL–1. The mixture of BSA and a-Lac was further applied to a gel filtration column to separate pure protein fractions. The yields of BSA and a-Lac, were about 2.3 and 1.1 mg mL–1, respectively, from the resolved peak. Brochier et al. (2008) demonstrated the feasibility of using a mixed-mode chromatography column for isolation of b-Lac from whey, without the necessity for pH or conductivity adjustment using a hexylamine mixed-mode column (HyperCel™, Pall BioSepra, Cergy, France). A smooth scalability from 2.5 to 10 mL column volume (CV) was achieved to extract all b-Lac from five CV of whey loaded into the column. Bound b-Lac was eluted at pH 4, with a purity estimated to be around 75%. Meanwhile, Schlatterer et al. (2004) used a ceramic hydroxyapatite column (Macro-Prep, BioRad, Munich, Germany) to isolate b-Lac from acid whey originating from the milk of healthy and mastitic cows. A single peak of b-Lac could be eluted at a sodium fluoride concentration of 0.6 M. Using
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460 Separation, extraction and concentration processes whey from a healthy cow, the yield of b-Lac was around 50–55%, with a purity of more than 96%. For mastitic whey, the yield of b-Lac was between 18 and 20%, with low purity, contaminated by immunoglobulins, BSA and lactoferrin. Membrane filtration In membrane filtration, b-Lac is normally recovered, together with BSA and immunoglobulin, as a retentate stream during ultrafiltration (UF). UF typically targets a-Lac as a permeant to produce very high purity. Bhattacharjee et al. (2006) used a two-stage UF membrane, followed by anion-exchange membrane chromatography, to produce a b-Lac fraction with 87.6% purity from whey protein concentrate. Bhushan and Etzel (2009) modified a UF membrane to have a positively charged membrane to separate two proteins of similar size, b-Lac and glycomacropeptide (GMP). b-Lac was retained by the charged membrane, whereas GMP selectivity was increased by over 600% compared with the uncharged membrane. Membrane chromatography Initial use of membrane chromatography for whey protein fractionation was carried out by Splitt et al. (1996) and Freitag et al. (1996). Splitt et al. (1996) demonstrated that chromatographic conditions were transferable from a cellulose- to a polymer-based membrane adsorber carrying the same functional groups for whey protein fractionation. In a two-step salt gradient, a-Lac was eluted at 0.1 M NaCl and a mixture of BSA and b-Lac were eluted at 0.5 M NaCl. By passing 5 L of feed (0.065 mg mL–1 a-Lac and b-Lac, 1 mg mL–1 BSA) through 1300 cm2 of total membrane adsorber area, 116 mg a-Lac was eluted in the first peak and 132 mg of a mixture of b-Lac and BSA was eluted in the second peak. Freitag et al. (1996) investigated the concept of mixed-mode interactions to bind all whey proteins in a single pass of whey through a membrane chromatography column. Two modules of MA Q15 and one module of MA S15 were connected in series and whey was passed through at pH 6. However, elution of the anion and cation modules was done separately, because a-Lac and IgG elute at the same salt concentration. MA Q15 produced a single peak containing a-Lac, BSA, b-Lac A and b-Lac B, whereas MA S15 gave a single peak of IgG. Recent studies by Saufi and Fee (2009) and Goodall et al. (2008) used anion exchange membrane chromatography to selectively bind b-Lac from whey. Goodall et al. (2008) observed that when the anionic membrane was saturated with whey, b-Lac could displace other bound protein from the membrane. This can produce a flow through fraction that is depleted in b-Lac, with concentrations of a-Lac and BSA doubled from their original concentrations in whey.
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Methods for purification of dairy nutraceuticals 461 16.3.2 Purification of a-lactalbumin Membrane filtration A novel variant of UF, known as high-performance tangential flow filtration (HPTFF), has been widely used in separation of a-Lac and b-Lac from whey. HPTFF exploits a number of different strategies to achieve high-resolution separations between proteins of similar size, by selecting specific conditions such as (Van Reis et al. 1997; 1999): ∑
proper choice of pH and ionic strength to maximize differences in the hydrodynamic volumes of the product and impurities, ∑ use of electrically charged membranes to enhance the retention of likecharged proteins, ∑ operation in the pressure-dependent regime to maximize selectivity, and ∑ use of a diafiltration mode to wash impurities through the membrane. A dramatic improvement in a-Lac permeation and selectivity can be achieved, as demonstrated by Cowan and Ritchie (2007), who obtained five times better selectivity of a-Lac after modifying a PES membrane with a positively charged group. Similar improvement was also shown by Lucas et al. (1998), who used an inorganic membrane coated with positively charged polyethyleneimine. The transmission of b-Lac was reduced to 1%, whereas a-Lac transmission was 10%, giving a selectivity close to 10. Chromatographic techniques Turhan and Etzel (2004) used an SP Sepharose Big Beads (Amersham Biosciences, Uppsala, Sweden) column to isolate a-Lac from lactic acid whey, achieving a purity of 93%. Gurgel et al. (2001) used a hexamer peptide ligand to bind a-Lac from whey protein isolates, attaining a purity of about 87%, with lactoferrin being the main impurity. Doultani et al. (2004) used a selective elution method to recover different fractions of whey protein bound onto a cation-exchange column at pH 4. Different elution solutions could be applied to produce: (1) a whey protein isolate (WPI), using 10 mM NaOH, (2) a-Lac, using 100 sodium acetate at pH 4.9 and WPI depleted in a-Lac, using 10 mM NaOH, (3) a-Lac, using 100 sodium acetate at pH 4.9, WPI depleted in a-Lac, using 50 mM sodium phosphate at pH 6.5, lactoperoxidase, using 0.35 M NaCl in 50 mM sodium phosphate at pH 6.5 and lactoferrin, using 1.20 M NaCl in 50 mM sodium phosphate at pH 6.5. According to the authors, this method offers the flexibility to switch between different protein fractions, day-to-day, depending on the market and customer demands.
Other techniques Chemical precipitation with sodium hexametaphosphate (Alomirah and Alli 2004) was used to recover a-Lac from whey. The yield of a-Lac was reported to be between 44 and 89%, with a-Lac purities between 86 and
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462 Separation, extraction and concentration processes 90% (Alomirah and Alli 2004). Tolkach et al. (2005) used selective thermal precipitation to isolate native a-Lac from whey protein concentrate. Before precipitation, the environmental parameters of WPC were optimized in terms of total protein, lactose and calcium content, and pH value. The purity of a-Lac achieved by selective thermal precipitation of optimized properties of WPC was 98%, with a recovery about 75%. Muller et al. (2003a, 2003b) used a two-step process to purify a-Lac from whey protein concentrates. In the first step, whey was filtered through a 30 kDa UF membrane with operating conditions that enhanced the ratio of a-Lac/b-Lac in the permeate stream by minimizing the passage of other whey proteins (Muller et al. 2003b). In the second step, two options were investigated, a second UF module or a selective precipitation route (Muller et al. 2003b). The precipitation route was more promising compared with UF, with the purity of a-Lac achieved in the range of 77–99% and yields of 46–83%, depending on the permeate properties from the first UF step. 16.3.3 Purification of bovine serum albumin Bovine serum albumin (BSA) is not commonly isolated as a specific product from whey but rather comes about as a by-product in a-Lac or b-Lac extraction methods. In UF, BSA is normally recovered in the retentate stream with b-Lac and immunoglobulins. In chromatographic separations, BSA typically binds to anion-exchange columns, but, by appropriate selection of elution buffers, BSA contamination in b-Lac or a-Lac fractions can be minimized.
16.4 Basic proteins in the dairy nutraceutical industry The most important basic proteins present in milk are lactoferrin (LF) and lactoperoxidase (LP). Lysozyme is also present in very minor quantities but it is not commercially viable to process this protein from milk. The concentration of LF, LP and lysozyme in milk varies from species to species, breed, stage of lactation, parturition, nutrition, udder health and season of lactation (Thomson et al. 2005). Below are some of the physicochemical characteristics of the basic whey proteins. 16.4.1 Lactoferrin Lactoferrin (LF) is a well-characterized iron-binding glycoprotein that belongs to the transferrin family, also known as lactotransferrin (LTF). The nonglycosylated form of LF has a molecular weight of 80 kDa (690 residues) and a pI of 8.6. LF is present in several secretory substances, including milk, tears and saliva (Masson et al. 1966). It is one of the major whey proteins in human milk, with a concentration of about 1.4–2.0 mg mL–1 but it is only a
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Methods for purification of dairy nutraceuticals 463 minor component in bovine milk, at only one-tenth of these concentrations (0.1–0.2 mg mL–1). Huge variations in lactoferrin concentrations (0.06–1.0 mg mL–1) of milk from individual cows has been monitored and reported in the literature (Indyk and Filonzi 2005). However, elevated levels of LF can be produced in bovine milk with the help of recombinant technologies (Thomassen et al. 2005). 16.4.2 Lactoperoxidase Lactoperoxidase (LP) was first identified and reported as an endogenous enzyme in milk by Arnold (1881). It has a single polypeptide chain with a molecular weight of approximately 78 kDa with 612 amino acid residues and a pI of 9.6 (Seifu et al. 2005). LP is one of the most abundant and heat-stable enzymes in the milk of many mammals and bovine milk contains about 20 times higher levels than that of human milk (Gothefors and Marklund 1975) and is the second most abundant enzyme in bovine milk. Concentrations of LP and its catalysing activity in milk depends on several factors, including feed, stage of lactation and the breed and health status of the animals (Fox and Kelly 2006). A typical concentration of LP was reported as 30–40 mg L–1 (Seifu et al. 2005) in bovine milk. The protein is less important than LF in terms of nutraceutical products but it is crucial in milk preservation and storage and binds to cation exchangers during isolation of LF. 16.4.3 Lysozyme Lysozyme is one of the most abundant enzymes present in human milk. Human lysozyme contains 130 amino acid residues and has a molecular mass of 14.7 kDa and a pI of 11.4. Lysozyme is present in a number of secretions, such as tears, saliva, urine, mucus and milk. Chicken egg white is the richest source of C-type lysozyme, having a concentration ranging between 3.4 and 5.8 g L–1 (Wilcox and Cole 1957; Sauter and Montoure 1972) and this protein is very closely related to a-lactalbumin in both sequence and structure. Although all mammals contain C-type lysozymes, they vary widely in terms of structure and physicochemical properties, such as folding/unfolding, structure, calcium binding, stability to heat and pH and pI. Among all mammalian species, ass (1428 mg L–1), mare (790–1330 mg L–1) and human milk (270–890 mg L–1) are the top three sources of lysozyme but bovine milk contains only a tiny amount (0.05–0.21 mg L–1) (Benkerroum 2008).
16.5 Purification technologies for basic whey proteins in the dairy nutraceutical industry Extraction and isolation of LF in both laboratory- and industrial-scale applications have been achieved using various processing technologies © Woodhead Publishing Limited, 2010
464 Separation, extraction and concentration processes such as ion-exchange, membrane-adsorption, size-exclusion, affinity and hydrophobic interaction-chromatography. The following section describes various separation technologies and their potential applications in the dairy industry. A number of chromatographic and membrane methods used for isolation of LF from various sources is given in Table 16.6. 16.5.1 Lactoferrin and lactoperoxidase purification Ion-exchange chromatography Amongst the basic proteins, LF has been extensively investigated in both laboratory and industrial-scale extraction technologies. However, in most cases LP has been co-eluted as a secondary product because of the similarity in their pI values. The most widely used extraction technology to isolate basic proteins has been cation exchange chromatography. Owing to the complex nature of milk, in general LF extraction is carried out by a sequence of individual processing methods, including casein precipitation, filtration and ionexchange chromatography. Cation-exchange matrices that might be routinely used in basic protein separations include phosphocellulose and carboxymethyl (CM) or sulfopropyl (SP) substituted cellulose, Sephadex and Sepharose. Among these cationic resins, CM-Sepharose (a weak cation exchanger) and SP-Sepharose (a strong cationic exchanger) resins (GE Healthcare, Sweden) are very well characterized for high throughput isolation of basic proteins from milk and whey. During early development, CM-Sepharose resin was the most widely used matrix for isolation of LF from acid and rennet whey at around pH 7.7, but in recent times most LF purification methods investigated have used SP-Sepharose strong cation exchange resins. SP-Sepharose Fast Flow (FF)™ or Big Beads™ and Streamline™-SP resins have been widely used for LF and LP extraction from milk and whey feeds in both lab and industrial-scale applications. Etzel et al. (2000) optimized a chromatographic process using SP-Sepharose Big Beads for isolation of LF from skim milk with >90% purities and 80% recoveries in a single-step packed-bed system. They also optimized Streamline-SP resin as an expanded-bed cation-exchange chromatography medium for isolation of LF and LP from skim milk with >80% recovery at a 200 cm h–1 flow rate. Billakanti and Fee (2009) characterized a cryogel monolith chromatography for extraction of minor proteins (LF) by cation-exchange chromatography from whole milk feeds at 550 cm h–1 in a single-step. Yield and purity of LF extracted using this process were >85 and 90%, respectively. Wu and Xu (2009) reported a purification process for isolation of LF and IgG from bovine colostrum using a serial cation–anion exchange chromatography system with 95 and 97% final purities, respectively. Lu et al. (2007) designed a productionscale technology for extraction of LF from bovine colostrum with final yields of 83% and purities of 94% using an ultrafiltration system coupled with a fast-flow strong cation-exchange chromatography column. Fee and Chand (2006) have investigated and optimized a process to isolate LF and LP from
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Table 16.6 Author and year
Purification examples for the basic whey proteins Configuration/material
Tu et al. 2002 Anti-LF immunoglobulins (IgYLF) immobilized on Sepharose 4B support Ye et al. 2000 Carboxymethyl cation-exchange chromatography
Target protein
Protein source Description of study/results
LF
Methods for purification of dairy nutraceuticals 465
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Colostral whey Final products have purification folds of 103–136 of initial concentration and recoveries of 82–99% Bovine They performed the adsorption process using 50 mM LFa and LFb colostrum phosphate buffer at pH 7.7 LFa was eluted in 0.38 M NaCl and LFb at 0.43 M NaCl buffer solutions. Also investigated variation in the LF concentrations of lactation time and species Li-Chan et al. Immobilized yolk antibody on a LF Milk and Binding capacity was a function of ligand density: at 1998 monoaldehyde-activated support cheese whey 9.2 mg mL–1 of IgYLF, binding capacity was 20% (mol %). This could be increased to 80% by using low ionic strength buffers Camperi et al. Red HE-3B dye affinity chromatography LF Rennet whey This process was achieved with 82% recovery and 98% 2000 purity Jyh-Ping and Microfiltration affinity purification LF and IgG Cheese whey Maximum binding capacity on heparin-Sepharose gel Cheng-Hsin was 124 mg mL–1. LF recovered in this process has 92% activity and 95% purity whereas IgG has 86% activity 1991 and 90% purity Korhonen Review Igs Bovine The latest developments and progress in separation 2004 colostrum, technologies for isolation of Igs in both small and largewhey and milk scale applications Stec et al. Isolation of polyclonal Igs by HPLC Igs Bovine serum Ammonium sulfate precipitation was performed as 2004 a first step of Igs isolation followed by HPLC. This method is applicable for preparative-scale applications DeSilva et al. Conference paper. Novel approaches Whey Dairy whey Recent developments in isolation and purification 2003 to meet the challenges in processing proteins and technologies of various proteins and peptides functional dairy components peptides
Author and year
Continued
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Target protein
Protein source Description of study/results
Xu et al. 2000 Separation of IgG and glycomacropeptides using a polystyrene anion exchanger IRA93 and Amicon YM100 membrane
IgG
Dairy whey
Akita and LiChan 1998
IgG1 and IgG2
Kim and LiChan 1998
Configuration/material
Specific antibody-immobilized immunoaffinity chromatography for isolation of bovine immunoglobulin G subclasses Single-step process for isolation of IgG using avidin-biotinylated IgY chromatography
Konecny et al. Purification of monospecific polyclonal 1994 antibodies from hyper immune bovine whey using immunoaffinity chromatography Labrou and Review Clonis 1994
IgG
Igs
Igs
Successful for concentration of IgG while removing all major whey protein contaminates. A continuous process using IRA 93 resin and 100 kD membranes, designed to produce enriched IgG GMP and WPI ingredients from rennet dairy whey Bovine milk, Gave >98% purity of IgG subclasses. Binding capacities colostrum and of the column were 27 and 38% (molar masses) for whey IgG1 and IgG2, respectively. Antibody columns were stable for more than a year with minimal antibody loss Cheddar Binding capacities achieved in this investigation were cheese whey 50–55% (w/w of IgG and IgY ligands). Bound IgG was eluted at pH 2.8 using glycine–HCl buffer with 99% purity Cheese whey
Various modes of affinity adsorption technologies used for isolation of Igs both in laboratory- and industrialscale applications
466 Separation, extraction and concentration processes
Table 16.6
Methods for purification of dairy nutraceuticals 467 whole milk at milking pH and temperatures using SP Sepharose Big Beads as a packed chromatographic resin. In their investigation, they successfully captured both LP and LF (total dynamic capacity of 49 mg mL–1 resin) from raw milk at 300 cm h–1 flow rate with minor leakage (<5%) in the flow through. Minor proteins captured directly from raw milk were expected to have improved activities and active yields. Andersson and Mattiasson (2006) used simulated moving bed (SMB) technology to isolate LF and LP from whey protein concentrate. SMB technology could be a potential application for large-scale purification processes because it attained 6.5 times the target protein concentration whilst consuming 4.3 times less buffer solution than conventional methods. Noel (2007) has described the largest chromatography column for industrial-scale isolation of LF using expanded-bed-adsorption (EBA) chromatography for processing 200 000 L of crude cheese whey per day. During their optimization process, they passed 85 000 L (three cycles) of crude cheese whey through 950 L of cation-exchange resin at a 900 cm h–1 linear velocity. Bound LF was eluted at mild alkaline conditions with a final product yield of 90–100% and very high purity (with a single band in SDSPAGE). The cation-exchange adsorbent used in this process had a dynamic binding capacity of 27–54 g L–1. Shiozawa et al. (2001) investigated EBA technology using Streamline-SP as a stationary phase for extraction of LF from skim milk with approximately 90% recovery and purity at a 150 cm h –1 linear flow rate. Uchida et al. (1996) obtained a patent for their successful industrial-scale isolation methods to purify LF, secretory components and LP from whey and milk with more than 80% purity of each component using gradient-elution protocols. LP was eluted first, followed by secretory components and then LF. Separated components were biologically active and could be utilized in pharmaceuticals, cosmetics, food supplements and drinks. Kussendrager et al. (1994) claimed a patent for developing an industrial-scale high-throughput chromatography process for isolation of LF and LP from skim milk and whey with more than 80% yields at high superficial velocities (>500 cm h–1). Fee and Chand (2005) reported on-farm capture of high-value minor milk proteins (LF and LP) from the raw milk of individual animals using a fully automated protein fractionation robotic system. This method was rapid, used a single step and was designed to minimize microbial contamination, while offering full product traceability to individual animals. Claycomb (2004) holds a patent for on-farm fractionation of high value milk proteins directly from raw milk using an automated milking system (AMS). Dionysius (1991) investigated a CM-Sephadex resin in a stirred tank system, mixing the resin with whey (cheese, rennet and acid) for 60 min at pH 7.0, and captured more than 80% of LP and 90% of LF, although the resin capacity changed with nature of whey. The bound LF was subsequently eluted by NaCl gradient elution. Heeboll-Nielsen et al. (2004) developed a superparamagnetic cation-
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468 Separation, extraction and concentration processes exchange chromatographic process for fractionation of bovine whey proteins including LF and LP with 28-fold purification over the starting material. The resin could be recovered magnetically in a high-gradient magnetic fishing system, allowing the resin size to be minimized for fast adsorption, while being recovered without the need for fine particle filtration, settling or centrifugation. Ulber et al. (2001) described a two-step membrane processing method for separation of LF from sweet whey, the first membrane removing lipids and other solids and the permeate then passing through a cation exchange membrane adsorber. Plate et al. (2006) used 2-m2 membrane modules for preparative-scale isolation of LF and LP from sweet whey with >90% recovery yields and purities. Chiu and Etzel (1997) described a method for extraction of LF and LP from cheese whey using microporous (3–5 mm) cationic membranes and achieved 50 and 70% recoveries, respectively. Mitchell et al. (1994) used cation-exchange membranes operated in either dead end or cross-flow configuration to isolate LF and LP from cheese whey. Affinity chromatography Affinity chromatography is the second most popular chromatography method for isolation of minor milk proteins from complex protein mixtures on a laboratory scale. Wolman et al. (2007) used Red HE-3B dye coupled with hollow fibers as a membrane affinity chromatography matrix for isolation of LF in a single step, achieving a higher adsorption capacity (111.0 mg mL–1) than the same ligand immobilized on agarose beads (9.3 mg mL–1 resins). They could achieve >94% purity and 91% recovery in the final product using whey and colostrum as feed materials. Chen et al. (2007) evaluated micrometre-sized monodisperse superparamagnetic polyglycidyl methacrylate (PGMA) particles coupled with heparin (PGMA-heparin) as an affinity method for isolation of LF from bovine whey in a single step. This magnetic affinity method resulted in a very high binding capacity of 164 mg g –1 resin and may hold promise as a fast process tool for industrial-scale purification of high-purity LF, although one must always consider the potential for leaching of toxic compounds from affinity ligands. Noppe et al. (2006; 2007) described immobilized affinity peptides as a new platform for rapid development of alternative affinity chromatography ligands for isolation of target proteins from crude feeds. In their method development, they used selected bacteriophages immobilized on a microporous monolith column for extraction of LF from complex fluids such as milk and blood serum. Bound LF was eluted with 1 M NaCl and fractions were recovered with >95% purity. Kawakami et al. (1987) used an immobilized LF-monoclonal antibody resin for a single-step isolation of LF from both human and bovine skim milk and colostrum, with very high purities and recoveries of >97%. Unlike LF, isolation of LP has attracted limited attention because of its low concentration and low commercial value. Shin et al. (2001) used
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Methods for purification of dairy nutraceuticals 469 immunoaffinity chromatography for selective isolation of LP from human whey but such processes have not been extensively studied for large-scale applications. Other chromatographic methods Size-exclusion (gel-filtration) chromatography is not commonly used as an early step in large-volume protein purification because linear velocities are low, but it is often used as a final purification step after other chromatography steps to enhance protein purity. For example, Al-Mashiki and Nakai (1987) applied size-exclusion chromatography for isolation of LF and immunoglobulins from different whey streams. Tomita et al. (2002) reported the possibility of LF and LP isolation from cheese whey or skim milk using a semi-large scale hydrophobic interaction chromatography method. Yoshida (1989) used hydrophobic interaction chromatography followed by DEAE ion exchange chromatography to isolate LF and LP from acid whey. Non-chromatographic methods Noh et al. (2005) selectively purified LF using a cationic surfactant as a micelle foaming agent, manipulating the protein behaviour by changing pH and salt concentration in the aqueous phase and surfactant concentration in the organic phase. LF was partitioned into the aqueous phase whereas all other proteins were solubilized into the organic phase. Fuda (2004) obtained approximately 25-fold enrichment of LF and LP from sweet whey using colloidal aphrons (CGAs), which are surfactant microbubbles generated by intense stirring of the anionic surfactant sodium bis-2ethylhexyl sulfosuccinate (AOT). Noel et al. (2001) investigated a low-cost foam fractionation process as the first step in separating LF from milk and achieved approximately 40% mass recovery at pH 10. Similarly, Saleh and Hossain (2001) developed a semi-batch foaming process to separate LF from a multicomponent mixture of BSA, a-Lac and LF by utilizing protein surface active properties. 16.5.2 Lysozyme purification Owing to its high isoelectric point (pI>11), lysozyme typically co-elutes with other basic proteins (LF and LP), but there have been few studies of the isolation of this protein. However, lysozyme has been successfully purified from milk or acid whey by a combination of affinity chromatography on heparin-Sepharose, followed by gel filtration on Sepharose 4B (Wang and Kloer 1984) or Sephadex-G50 (Boesman-Finkelstein and Finkelstein 1982). Duhaiman (1988) used heparin–Sepharose 4B, Sephadex G-75 and hydroxyapatite for chromatographic extraction of lysozyme from camel milk. Recently, using recombinant technology, human lysozyme was expressed at high concentrations (0.5% dry weight) in transgenic rice seed (Huang et
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470 Separation, extraction and concentration processes al. 2002). Wilken and Nikolov (2006) found cation-exchange purification of lysozyme to be optimal at pH 4.5.
16.6 Immunoglobulins in the dairy nutraceutical industry IgG is an important component of bovine milk and comprises the major protein in bovine colostrum. Consumer acceptance of value-added whey protein products that involve Igs began in Asia and migrated to Europe and the US, targeting sports nutrition, infant formulae, dietary supplements and physiologically functional foods (Gapper et al. 2007). Immunoglobulins are g-globulin proteins that are found in blood and other body fluids of human, bovine and all other lactating species. These are used in the immune system to identify and neutralize bacteria, viruses and other antigens. Human immunoglobulins are broadly classified into five different classes: IgA, IgD, IgE, IgG, and IgM, whereas bovine milk and colostrum mainly contain immunoglobulins IgG, IgM and IgA. All immunoglobulin classes share a basic structural unit that comprises four polypeptide chains with two identical heavy (H) chains and two identical light (L) chains, held together with disulfide bonds. Molecular masses of the heavy- and light-chain peptides are approximately 50–70 kDa and 25 kDa, respectively. 16.6.1 Immunoglobulin G Immunoglobulin G is the most abundant class of antibody in the colostrum and milk of several mammalian species, comprising 80–90% of total antibodies. The primary structure of IgG contains two heavy (each of 450–550 amino acids) and two light chains (each of 211–217 amino acids) with a total molecular mass of approximately 150–160 kDa and a pI of 5.5–8.3. Commonly, IgG presents in two-sub classes such as IgG1 (MW 160 kDa, pI 5.5–6.8) being the richest class, at 15–180 g L–1, and IgG2 (MW 150 kDa, pI 7.5–8.3), at 1–3 g L–1. 16.6.2 Immunoglobulin A Similar to other immunoglobulins, the monomeric structure of IgA also contains two heavy and two light chains but it occurs as a monomer or a dimer, the latter joined by a J-chain and a secretory component. This complex structure is called secretory IgA (sIgA) and has a molecular weight of 380–417 kDa. IgA is present in blood, milk and several mucosal surfaces of lungs and gastrointestinal tracts. This antibody acts as a primary defence system against several pathogens. The IgA concentration in bovine colostrum ranges from 1 to 5 g L–1, whereas bovine milk contains only 0.05 to 1 g L–1.
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Methods for purification of dairy nutraceuticals 471 16.6.3 Immunoglobulin M The monomeric structure of IgM is similar to IgG but IgM consist of five subunits, linked together to form a pentameric structure in a circular mode by disulfide bonds and a J-chain. The molecular mass of IgM is approximately 900–1000 kDa. In bovine milk, the IgM concentration ranges from 0.04 to 1.0 g L–1, whereas bovine colostrum contains 5 to 10 times greater amounts.
16.7 Purification technologies for immunoglobulins in the dairy nutraceutical industry Immunoglobulins from bovine milk have applications as supplements in infant formulae, hyperimmune foods, functional foods, nutraceutical and pharmaceutical products. Although there are many processing methods available for separation of immunoglobulins from various sources, including colostrum and milk, appropriate technologies for large-scale productions are still lacking. Traditionally, isolation of Igs from colostral whey has been achieved by precipitation with either ammonium sulfate or ethanol, followed by chromatography. Although such methods yield rather pure IgG fractions, most are feasible only in small-scale applications. In particular, they are not appropriate for large-scale production of Igs from bovine colostrum or milk (El-Loly 2007). Because milk is a complex fluid and contains high amounts of fat, it is difficult to isolate Igs using conventional methods. However, rapid developments in separation technologies and particularly in the application of membrane separation methods have made industrial-scale isolation of immunoglobulins from various streams, including whey, colostrum and recombinant cell culture supernatants a possibility. 16.7.1 Ion exchange chromatography Use of ion-exchange chromatography to isolate immunoglobulins from dairy feeds is limited but it is used as an intermediate or final purification step in combination with other chromatographic or membrane processing methods. For example, Wu and Xu (2009) applied cation- and anion-exchange chromatography in series for isolation of IgG and LF from bovine colostral whey. Similarly, Pessela et al. (2006) investigated a simple method for isolation of Igs from whey protein concentrate with >80% recovery and reasonable purity using a combination of DEAE–agarose anion-exchange chromatography followed by low-substitution aminated adsorbents. In the first step, they could remove the main contaminant, BSA, and during the second step, Igs were selectively bound on the aminated resin with minor contamination by a-Lac and b-Lac. Noel (2007) developed a second-generation, robust EBA method for
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472 Separation, extraction and concentration processes isolation of Igs from cheese whey, which involved a mixed-mode chemical ligand that has selectivity towards IgG over other whey proteins. The feed whey solution was heated to 50 °C to maintain pasteurization, reduce the viscosity of the solution and maximize flow rate. They could thus load the column at greater than 3000 cm h–1 flow rate without reducing its adsorption capacity. Using a 880-L EBA column, they could produce 13 kg of IgG per day with >90% of final product yields. The eluent from this column was passed through a 160 L anion exchange column for further purification. 16.7.2 Membrane chromatography Since membrane technology was introduced to the dairy industry in the 1970s, many membrane processes have been developed for isolation of Igs from whey and milk using either a single-step process or a combination of several individual steps. For example, UF has been employed either alone or in combination with ion-exchange or gel-filtration chromatography for largescale fractionation of Igs from cheese or colostral whey streams (Abraham 1988; Korhonen 2004; Syväoja et al. 1994). Al-Mashikhi et al. (1988) successfully isolated IgG from acid and cheddar whey using UF, combined with immobilized metal chelate affinity chromatography, obtaining purities of 77 and 53%, respectively. Korhonen (1998) used a multi-step process with UF, microfiltration and reverse osmosis, followed by a cation-exchange resin to achieve 40–75% enrichment of Igs from colostral whey. 16.7.3 Affinity chromatography Affinity adsorption methods have been widely used for isolation of Igs from colostral or cheese whey. Protein-A and Protein-G affinity media are well characterized and used for fractionation of Igs, particularly IgG. Many pilot- and industrial-scale methods using membrane technologies and affinity chromatography have been developed for extraction of Igs from whey and colostrum either alone or in combination with other chromatography techniques (Kochan et al. 1996; Mukkur and Froese 1971; Schmerr et al. 1985). Labrou and Clonis (1994) reviewed various affinity-based purification technologies for Igs. Human and secretory IgA molecules have been isolated by precipitation, ion-exchange and affinity chromatography from human serum (Kondoh et al. 1987; Monteiro et al. 1985; Roque-Barreira et al. 1986) and human milk (Khayam-Bashi et al. 1977). Hutchens et al. (1990) developed a salt-promoted thiophilic adsorption method for simultaneous isolation of all bovine Igs from colostral whey. Mukker and Froese (1971) described isolation of bovine IgM from colostral whey using a combination of molecular sieve and DEAE–cellulose chromatography. Su and Chiang (2003) described a simple non-chromatographic method for isolation of Igs from other whey proteins. They used reverse micelles
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Methods for purification of dairy nutraceuticals 473 to extract more than 90% of IgG into an aqueous phase, with 90% purity, whereas all other whey proteins were partitioned into the reverse micelle phase. 16.7.4 Protein-A mimetic (PAM) affinity peptides Protein A and protein G affinity chromatography can achieve capacities of up to 30–40 mg IgG per mL of resin under optimal conditions. However, these affinity ligands suffer from several inherent drawbacks such as high cost, their biological origin, possible contaminants, low stability towards sanitizing agents and the fact that their selectivity is limited to IgG. Recently, with the help of combinatorial chemistry and modern analytical tools, researchers synthesized protein A mimetic (PAM) peptide ligands (D’Agostino et al. 2008; Fassina 1994; Newcombe et al. 2005; Ulbricht and Yang 2005; Yang et al. 2005) for isolation of human immunoglobulins from mammalian cell cultures. Compared with protein A or G, these synthetic peptide ligands have several benefits (D’Agostino et al. 2008), including high stability towards chemical and biological reagents, low production cost and the absence of biological contaminants. Moreover, these synthetic affinity ligands may also have selectivity towards other classes of immunoglobulins (IgG, IgE, IgM and IgA) (Fassina et al. 1998; Yang et al. 2005), which are not recognized by conventional protein A or G ligands. Several synthetic peptides have been proposed as potential replacements for protein A in affinity chromatography (Ulbricht and Yang 2005). However, the lack of selectivity for antibodies from various classes has limited their widespread use and some of these ligands are not compatible with the retention of antibody activities in proposed buffer solutions. However, more recently, Yang et al. (2005) synthesized a hexamer peptide that has a wide-ranging selectivity towards various classes of immunoglobulins at near-neutral pH. By using a gradient elution with 0.2 M sodium acetate buffer, they were able to isolate various classes of immunoglobulins from human serum (Yang et al. 2005). They have also reported the specificity of this ligand towards immunoglobulins of human, bovine and several other species. Although most of this work has been carried out on human immunoglobulins, there has been recent interest in applying this PAM peptide affinity chromatography for isolation of bovine immunoglobulins from early colostral milk, hyperimmunized cows and transgenic animals.
16.8 Future trends In the past few decades, the dairy industry has moved from being solely a commodity food-based industry producing significant waste (whey) volumes to one that has increasingly sought added value and the minimization of waste. The nature of dairy products has expanded from simple nutritional © Woodhead Publishing Limited, 2010
474 Separation, extraction and concentration processes foods to separate fractions sold as individual ingredients or as blends to meet the specific functional and nutritional requirements of secondary producers. This trend will continue, as milk volumes increase to meet the new demand for dairy products from the developing world and greater bulk production efficiencies drive companies to find new products to maintain their competitive advantages. Nutraceutical products, with implicit claims of health benefits, are now commonplace and there are likely to be more such products in the future, with some fractions in the future perhaps even meeting the more stringent proof of efficacy requirements of pharmaceuticals for highly targeted applications, for example around neonatal gut development, anticancer agents or immune system enhancers. The commercial recovery of very minor fractions with useful bioactivities, such as oligosaccharides (including those incorporating sialic acid), the myriad proteins that make up the milk fat globule membrane, and the complex lipids therein (phosphatidylcholine, phosphatidylethanolamine, phosphatidylserine, phosphatidylinositol, sphingomyelin and glycolipids such as gangliosides) will be a challenge. Recovery of such fractions may require specific affinity techniques and the challenge will be to use techniques to extract minor proteins economically on a large scale, while maintaining the safety and integrity of the residual milk or whey for further processing i.e. without contaminating the flow-through stream with leachates. On-farm extraction of minor products from milk may offer advantages in the production of proteins and other bioactives that are heat-labile and are currently largely lost during the pasteurization required inside the factory gate. On-farm extraction may also offer direct benefit to farmers who select for certain genetic traits or apply unique feeding regimes to enhance the production of specific high-value products in their herds. In summary, higher volumes of milk production and increased global competition in the dairy industry will probably drive further growth in niche bioactive products over the next few decades. High-volume, economically efficient separation technologies with high specificity and compatibility with food production will be required to support this growth.
16.9 References Abraham, G. B. (1988). Process for preparing antibodies against E. coli K-99 antigen from bovine milk, US patent 4784850. Akita, E. M. and Li-Chan, E. C. Y. (1998). ‘Isolation of bovine immunoglobulin G subclasses from milk, colostrum, and whey using immobilized egg yolk antibodies’. Journal of Dairy Science 81(1): 54–63. Al-Mashikhi, S. A., Li-Chan, E. and Nakai, S. (1988). ‘Separation of immunoglobulins and lactoferrin from cheese whey by chelating chromatography’. Journal of Dairy Science 71(7): 1747–1755. Al-Mashikhi, S. A. and Nakai, S. (1987). ‘Isolation of bovine immunoglobulins and lactoferrin from whey proteins by gel filtration techniques’. Journal of Dairy Science 70(12): 2486–2492.
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480 Separation, extraction and concentration processes Brazilian Journal of Medical and Biological Research 19(2): 149–157. Saleh, Z. S. and Hossain, M. M. (2001). ‘A study of the separation of proteins from multicomponent mixtures by a semi-batch foaming process’. Chemical Engineering and Processing 40(4): 371–378. Saufi, S. M. and Fee, C. J. (2009). ‘Fractionation of b-lactoglobulin from whey by mixed matrix membrane ion exchange chromatography’. Biotechnology and Bioengineering 103(1): 138–147. Sauter, E. A. and Montoure, J. E. (1972). ‘The relationship of lysozyme content of egg white to volume and stability of foams’. Journal of Food Science 37(6): 918–920. Schlatterer, B., Baeker, R. and Schlatterer, K. (2004). ‘Improved purification of b-lactoglobulin from acid whey by means of ceramic hydroxyapatite chromatography with sodium fluoride as a displacer’. Journal of Chromatography B: Analytical Technologies in the Biomedical and Life Sciences 807(2): 223–228. Schmerr, M. J. F., Patterson, J. M., Van Der Maaten, M. J. and Miller, J. M. (1985). ‘Conditions for binding bovine IgG1 to protein A-sepharose’. Molecular Immunology 22(5): 613–616. Seifu, E., Buys, E. M. and Donkin, E. F. (2005). ‘Significance of the lactoperoxidase system in the dairy industry and its potential applications: A review’. Trends in Food Science and Technology 16(4): 137–154. Semo, E., Kesselman, E., Danino, D. and Livney, Y. D. (2006). Casein micelle as a natural nano-capsular vehicle for nutraceuticals. Food Colloids 2006, April 23–26, Montreux, Switzerland. Severin, S. and Xia, W. S. (2005). ‘Milk biologically active components as nutraceuticals: Review’. Critical Reviews in Food Science and Nutrition 45(7–8): 645–656. Shea, A. P., Crofcheck, C. L., Payne, F. A. and Xiong, Y. L. (2009). ‘Foam fractionation of a-lactalbumin and b-lactoglobulin from a whey solution’. Asia-Pacific Journal of Chemical Engineering 4(2): 191–203. Shin, K., Hayasawa, H. and Lonnerdal, B. (2001). ‘Purification and quantification of lactoperoxidase in human milk with use of immunoadsorbents with antibodies against recombinant human lactoperoxidase’. American Journal of Clinical Nutrition 73(5): 984–989. Shiozawa, M., Okabe, H., Nakagawa, Y., Morita, H. and Uchida, T. (2001). ‘Purification of lactoferrin by expanded-bed column chromatography’. Kagaku Kogaku Ronbunshu 27(2): 147–148. Spitsberg, V. L. (2005). ‘Bovine milk fat globule membrane as a potential nutraceutical’. Journal of Dairy Science 88(7): 2289–2294. Splitt, H., Mackenstedt, I. and Freitag, R. (1996). ‘Preparative membrane adsorber chromatography for the isolation of cow milk components’. Journal of Chromatography A 729(1–2): 87–97. Stec, J., Bicka, L. and Kuźmak, J. (2004). ‘Isolation and purification of polyclonal IgG antibodies from bovine serum by high performance liquid chromatography’. Bulletin of the Veterinary Institute in Pulawy 48(3): 321–327. Su, C. K. and Chiang, B. H. (2003). ‘Extraction of immunoglobulin-G from colostral whey by reverse micelles’. Journal of Dairy Science 86(5): 1639–1645. Suutari, T. J., Valkonen, K. H., Karttunen, T. J., Ehn, B. M., Ekstrand, B., Bengtsson, U., Virtanen, V., Nieminen, M. and Kokkonen, J. (2006). ‘IgE cross reactivity between reindeer and bovine milk b-lactoglobulins in cow’s milk allergic patients’. Journal of Investigational Allergology and Clinical Immunology 16(5): 296–302. Syväoja, E.-L., Ahola-Luttila, H. K., Kalsta, H., Matilainen, M. H., Laakso, S., Husu, J. R. and Kosunen, T. U. (1994). ‘Concentration of campylobacter-specific antibodies in the colostrum of immunized cows’. Milchwissenschaft 49: 27–31. Tellez, C. M. and Cole, K. D. (2000). ‘Preparative electrochromatography of proteins in various types of porous media’. Electrophoresis 21(5): 1001–1009. Thomassen, E. A. J., Van Veen, H. A., Van Berkel, P. H. C., Nuijens, J. H. and Abrahams,
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482 Separation, extraction and concentration processes Yang, H., Gurgel, P. V. and Carbonell, R. G. (2005). ‘Hexamer peptide affinity resins that bind the Fc region of human immunoglobulin G’. Journal of Peptide Research 66(Suppl. 1): 120–137. Ye, X., Yoshida, S. and Ng, T. B. (2000). ‘Isolation of lactoperoxidase, lactoferrin, a-lactalbumin, b-lactoglobulin B and b-lactoglobulin A from bovine rennet whey using ion exchange chromatography’. International Journal of Biochemistry and Cell Biology 32(11–12): 1143–1150. Yoshida, S. (1989). ‘Preparation of lactoferrin by hydrophobic interaction chromatography from milk acid whey’. J. Dairy Sci. 72: 1446–1450. Zydney, A. L. (1998). ‘Protein separations using membrane filtration: new opportunities for whey fractionation’. International Dairy Journal 8(3): 243–250.
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17 Methods of concentration and purification of omega-3 fatty acids S. P. J. Namal Senanayake, Danisco USA, Inc., USA
Abstract: An overview is presented of the various methodologies used for producing highly purified omega-3 fatty acids from natural source materials. Omega-3 fatty acids derived from fish, krill and microalgae, consisting of eicosapentaenoic acid (EPA; 20:5n-3), and docosahexaenoic acid (DHA; 22:6n-3), have beneficial effects in the prevention and management of cardiovascular disease and other chronic disorders. Production of high-purity omega-3 fatty acids is increasingly important in both the nutraceutical and pharmaceutical industries. The physical, chemical and enzymatic methods used include urea adduction, chromatography, low-temperature fractional crystallization, supercritical fluid extraction and distillation. Key words: omega-3 fatty acids, urea adduction, chromatography, crystallization, supercritical fluid extraction, distillation, enzymatic methods.
17.1 Introduction Omega-3 fatty acids are essential fatty acids that have diverse biological effects in human health and disease. They are essential to human health but cannot be manufactured by the body. For this reason, omega-3 fatty acids must be obtained from food. Omega-3 polyunsaturated fatty acids (PUFAs) can be found in fish, such as salmon, tuna, and halibut, other marine organisms such as algae and krill, certain plants, and nut oils. The heart-health benefits of the omega-3 fatty acids are well known. Omega-3 fatty acids are considered to have beneficial effects in the prevention of cardiovascular disease, inflammation, hypertension as well as other chronic disorders (Kris-Etherton et al., 2001; Lands, 2003; Senanayake, 2000; Senanayake and Shahidi, 2000b; Shahidi and Senanayake, 2006; Yokoyama et al. 2007). In 2004, the US Food and © Woodhead Publishing Limited, 2010
484 Separation, extraction and concentration processes Drug Administration gave ‘qualified health claim’ status to eicosapentaenoic acid (EPA; 20:5n-3) and docosahexaenoic acid (DHA; 22:6n-3) found in fish, stating that ‘supportive but not conclusive research shows that consumption of EPA and DHA omega-3 fatty acids may reduce the risk of coronary heart disease’ (United States Food and Drug Administration, 2004). For clinical and nutritional applications, the natural sources of omega-3 fatty acids, as such, may not provide the required amounts of these fatty acids and hence concentration and purification of omega-3 fatty acids may be necessary. A concentrated source of these fatty acids is desired to achieve sustainable benefits. Highly purified omega-3 fatty acids may be produced in the free fatty acid, simple alkyl ester and triacylglycerol forms. To achieve this, physical, chemical and enzymatic techniques may be employed. Methods traditionally employed for the concentration and purification of omega-3 fatty acids include urea adduction, chromatographic methods, low-temperature fractional crystallization, supercritical fluid extraction, distillation, and enzymatic and integrated methods.
17.2 Urea adduction in the concentration and purification of omega-3 fatty acids Urea adduction is one of the most efficient and simplest techniques for concentration and purification of omega-3 fatty acids from natural sources (Senanayake, 2000). The formation of complexes between urea and straightchain saturated fatty acids is a well established and potentially valuable separation technique for fractionation of free fatty acids or esters (Shahidi and Senanayake, 2006). Initially, the triacylglycerols (TAGs) of oil are hydrolyzed into their constituent fatty acids via alkaline hydrolysis using alcoholic KOH or NaOH. The resultant free fatty acids (FFAs) are then mixed with an ethanolic solution of urea for complex formation. Urea molecules readily form solid-phase complexes with saturated and monounsaturated fatty acids and crystallize out on cooling and may be removed by filtration. The liquid fraction is highly enriched with omega-3 fatty acids. In general, the crystallization temperature can be ranged from ambient to –20 °C. Since the process involves relatively mild operating conditions and chemicals (free fatty acids, urea, ethanol, water) that are generally recognized as safe (GRAS) by the US Food and Drug Administration, it can be labeled as eco-friendly. This process can be used as a preliminary fractionation step in conjunction with other purification methods, such as low-temperature fractional crystallization, enzyme-catalyzed methods, and molecular distillation, resulting in a highpurity free-fatty-acid omega-3 product. Urea-based fractionation of fatty acids has been extensively used for enriching omega-3 fatty acids in marine oils (Hayes et al., 2000). Urea adduction of fatty acid ethyl esters of squid (Illex argentinus) visceral oil,
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using cyclohexane and methanol, can increase the EPA content from 11.8 to 28.2% and DHA content from 14.9 to 35.6% (Hwang and Liang, 2001). Haagsma et al. (1982) described a urea adduction method for enriching the EPA and DHA levels of cod liver oil from 12 to 28 and 11 to 45%, respectively. Wanasundara and Shahidi (1999) used an experimental design to optimize the conditions on a laboratory scale that led to the maximum concentration of EPA and DHA from seal blubber oil, using the urea-based fractionation. The authors obtained 88.2% of total omega-3 fatty acids at an urea/fatty acid ratio of 4.5, a crystallization time of 24 h and a crystallization temperature of –10 °C. Senanayake and Shahidi (2000a) concentrated and purified DHA from the oil extracted from the microalgae Crypthecodinium cohnii and reported a DHA enrichment from 47.4 to 97.1% with a process yield of 32.5% of the mass of the original algal oil. The fatty acid composition of algal oil and its DHA concentrate obtained by urea adduction are reported in Table 17.1 (Senanayake and Shahidi, 2000a). In another study, mackerel processing waste comprising skins, viscera, and muscle tissue was evaluated by Zuta et al. (2003) for concentrating omega-3 fatty acids by urea adduction. Fish oil was extracted using either chloroform/methanol (2:1 vol/vol) or hexane/isopropanol (3:2 vol/vol). The oil yield and iodine value (which measures the degree of unsaturation in fats and oils) were determined for fresh fish oil extracts. Omega-3 fatty acid concentrates were prepared from saponified fish oil via urea adduction. The mean oil yields were 9.18, 9.2, and 38.1% for viscera, muscle, and skin, respectively. The mean baseline iodine value was 134, which increased to 296 after urea adduction. Hence, it was possible to concentrate omega-3 fatty acids from mackerel processing waste. The type of tissue used did not affect the amount of omega-3 fatty acids concentrated. Mackerel skin was most desirable because of its high oil content. Hayes (2006) investigated the effect of cooling rate on the degree of removal of saturated acyl groups from FFAs derived from canola oil and Table 17.1 Fatty acid composition of algal oil and its DHA concentrate after urea adduction (adapted from Senanayake and Shahidi, 2000a) Fatty acids 10:0 12:0 14:0 16:0 16:1 18:1w-9 18:2 w-6 22:5 w-3 22:6 w-3 Iodine value, calculated
Algal oil (%) 0.6 1.1 15.0 9.0 2.2 19.0 1.0 0.5 47.4 234
DHA concentrate (%) 0.5 0.5 0.1 0.0 0.3 0.2 0.7 0.4 97.1 437
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486 Separation, extraction and concentration processes the isolation of di- and polyunsaturated acyl groups from FFAs derived from vegetable and fish oil during urea adduction. Traditionally, slow cooling has been used (–1 °C min–1). A more rapid cooling rate (–47 °C min–1) produced urea inclusion compounds (crystals) of similar morphology and thermodynamic properties, but of a size an order of magnitude smaller than the urea inclusion compounds formed during slow cooling. Fractionations used only renewable materials (urea, FFAs, and 95% ethanol as solvent) and benign operating conditions (ambient pressure, 25–75 °C, and neutral pH). When the recovery of FFAs was relatively high (>60%), the selectivity of urea adduction toward the inclusion of saturated fatty acids and against PUFAs was not affected by the cooling rate. In contrast, when the FFAs recovery was low, representing those instances where an increase in the purity of the PUFAs is a more important economic goal, a slower cooling rate resulted in a significantly greater discrimination against PUFA groups, hence to a FFA product with a measurably greater purity.
17.3 Chromatographic methods for the concentration and purification of omega-3 fatty acids Another method for concentration and purification of omega-3 fatty acids is the use of chromatography. High-performance liquid chromatography, silver resin chromatography and supercritical-fluid chromatography have been used for concentration of omega-3 fatty acids from natural sources. Hayashi and Kishimura (1993) isolated 63–74% pure DHA from skipjack tuna eye orbital oil by stepwise elution with hexane, diethyl ether/hexane and diethyl ether on a silicic acid column. Teshima et al. (1978) employed a silver nitrate-impregnated silica gel column to separate DHA and EPA from squid-liver oil after forming methyl esters. They were able to isolate 95–98% DHA and 85–96% EPA with yields of 48 and 39%, respectively. Guil-Guerrero and Belarbi (2001) purified EPA and DHA from cod liver oil using a silver nitrate-impregnated silica gel column. The oil was saponified and treated with urea, and the non-complexed fatty acids were then converted to methyl esters before chromatography. The column was washed with a sequence of mobile phases. A 64% recovery of DHA with 100% purity was obtained. The recovery of EPA was 29.6%, with a final purity of 90.6%. Perretti et al. (2007) investigated the fractionation of fish oil fatty acid ethyl esters with the aim of obtaining a lipid fraction enriched in omega-3 fatty acids and with a suitable EPA/DHA ratio. They reported the possibility of modifying the original fatty acid ethyl ester concentrations by optimizing the extraction conditions in terms of pressure, temperature, and supercritical CO2 flow rate: 2-h runs, pressures of 100, 140, 150, and 300 bar, and liquid CO2 flow rates of 2.5, 3.5, 5, and 10 kg h–1, at 40, 50, and 60 °C, in the three sections of the column starting from the bottom, respectively. They
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stated that supercritical fluid fractionation appears to be a useful processing technique for changing the composition of lipids in order to obtain high added-value functional products. A preparative reversed-phase high-performance liquid chromatography method with gradient elution using acetonitrile–chloroform and evaporative light-scattering detection was used by Mansour (2005) to purify milligram quantities of microalgal PUFAs, separated as methyl esters. PUFA-methyl esters purified included methyl esters of DHA, EPA and the unusual very-longchain highly unsaturated fatty acid, octacosaoctaenoic acid (28:8n-3) from the marine dinoflagellate, Scrippsiella sp. Other fatty acids purified from various microalgae using this method to greater than 95% purity included 16:3(n-4), 16:4(n-3), 16:4(n-1) and 18:5(n-3). The number of injections required was variable and depended on the abundance of the desired PUFA-methyl esters, and resolution from closely eluting PUFA-methyl esters, which determined the maximum loading. The purity of these fatty acids was determined by electron-impact gas chromatography–mass spectrometry (EI GC–MS) and the chain length and location of double bonds were determined by EI GC– MS of 4,4-dimethyloxazoline derivatives formed using a low-temperature method. The advantages over silver-ion HPLC for purifying PUFA-methyl esters are that separation occurs according to chain length as well as degree of unsaturation enabling separation of PUFA-methyl esters with the same degree of unsaturation but different chain length. In addition, PUFA-methyl esters were not strongly adsorbed, but eluted earlier than their more saturated corresponding PUFA-methyl esters of the same chain length. This method is robust, simple, and requires only a short re-equilibration time. Centrifugal partition chromatography (CPC) is a separation method based on the liquid partition of compounds. Owing to the high centrifugal field force, one phase stays in the rotor (the liquid stationary phase), the other one is the mobile phase as in classical liquid chromatography. CPC has been used for isolation and purification of PUFAs from various natural sources. Several important advantages of centrifugal partition chromatography are reported in Table 17.2. A CPC method to purify DHA from microalgal oil on a laboratory scale, has been developed by Wanasundara and Fedec (2002). The starting microalgal oil contained 39.7% DHA and 15.2% DPA (n-6), along with several other fatty acids. The mobile phase was hexane–methanol–water in a normal phase-ascending mode. A good separation was achieved and recoveries of 84.6% DHA and 84.9% DPA were obtained. Bousquet and Goffic (1995) have examined CPC separation of microalgal oil-based DHA and EPA and were able to isolate pure DHA and EPA from this oil with excellent yields. The first separation used heptane as the stationary phase and aqueous 3% acetonitrile as the mobile phase. The minor fatty acids were eluted, leaving a mixture of four major PUFAs. This mixture was subjected to a second separation, using heptane as a stationary phase and aqueous methanol as the mobile phase. Under these conditions, it was possible to isolate pure DHA and EPA with high yields. High-speed countercurrent chromatography was used by Du et
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488 Separation, extraction and concentration processes Table 17.2 Advantages of centrifugal partition chromatography Advantages Cost
Solid supports, which in many cases are costly, are not required but instead a solvent system is used Recovery Elimination of solid supports avoids problems associated with irreversible retention of highly retentive sample components. Almost 100% recovery of the compounds is guaranteed Quantity Compared with conventional liquid chromatography, the volume ratio of the stationary phase to the total column (rotor) volume is greater. Hence, large quantities of materials can be retained in the stationary phase Speed The stationary phase is retained by centrifugal force, enabling the mobile phase to be pumped at high speeds through the apparatus resulting in reduced separation times Stability Mild operating conditions are used. Therefore, decomposition and oxidation of PUFAs are virtually nonexistent under these conditions Versatility Any two-phase solvent system can be used, including many prepared from nontoxic and commonly available solvents
al. (1996) to separate DHA and EPA ethyl esters in a mixture containing 39.3 and 56.4%, respectively, with hexane–dichloromethane–acetonitrile (5:1:4 v/v/v) as mobile phase. However, the separation was not successful with this solvent system because of the low partition coefficients of the esters under the experimental conditions used. Murayama et al. (1988) separated a mixture of stearic, oleic, linoleic and a-linolenic (18:3n-3) acid ethyl esters because their partition coefficients were distributed over a wide range in the mobile phase comprising hexane–acetonitrile (1:1 v/v). The ethyl esters of linoleic acid and a-linolenic acids were successfully separated during the first normal ascending elution whereas the ethyl ester of oleic acid was isolated by switching the elution mode to descending. Despite the developments in chromatographic techniques to concentrate and purify omega-3 oils, the use of very large volumes of solvents, loss of column resolution after repeated use, and potential product solvent residues are likely to hinder scale-up to production scale volumes.
17.4 Low-temperature fractional crystallization for the concentration and purification of omega-3 fatty acids Low-temperature fractional crystallization is one of the simplest methods employed for production of purified omega-3 fatty acids. This process takes advantage of the existing differences in the melting points of different fatty acids as neat oils or in different solvent systems. The more saturated fatty acids have higher melting points and crystallize out of the mixtures
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leaving behind the more unsaturated fatty acids (Senanayake, 2000). The melting points of fatty acids are dependent on their degree of unsaturation. For example, EPA and DHA melt at –54 and –44.5 °C compared with 13.4 and 69.6 °C for 18:1 and 18:0, respectively (Merck Index, 1983). As the temperature of a mixture of saturated and unsaturated fatty acids decreases, the saturated fatty acids, having a higher melting point, start to crystallize out first and the liquid phase becomes enriched in the unsaturated fatty acids (Shahidi and Senanayake, 2006). However, as the number and type of fatty acid components in the mixture increases, the crystallization process becomes more complex and repeated crystallization and separation of fractions must be carried out to obtain purified fractions. For fish oils, not only is there a very wide spectrum of fatty acids but the fatty acids exist, not in the FFA form, but esterified in TAGs. However, the principle of low-temperature fractional crystallization can still be applied to marine oils to partially concentrate TAGs rich in omega-3 PUFAs (Shahidi and Wanasundara, 1998). Seal blubber oil in the TAG and FFA forms were subjected to low-temperature fractional crystallization using solvents such as hexane and acetone in order to obtain omega-3 fatty acid concentrates. Preparation of omega-3 fatty acid concentrates from seal blubber oil by low-temperature crystallization was also reported by Wanasundara (1996), who subjected seal blubber oil in the TAG or FFA form to solvent-based fractionation, using hexane and acetone as solvents, at different temperatures. The content of omega-3 fatty acids in the non-crystalline fraction was increased with decreasing the crystallization temperature. Under all temperature conditions, acetone gave rise to the highest concentration of total omega-3 fatty acids. Low-temperature solvent crystallization of seal blubber oil, in the FFA form, at –60 and –70 °C in hexane, resulted in total omega-3 fatty acid content of up to 58.3 and 66.7%, respectively. However, the content of total omega-3 fatty acids in acetone increased up to 56.7 and 46.8%, respectively. In another study, Han et al. (1987) found that alkali salts of less unsaturated fatty acids crystallize more rapidly than those of highly unsaturated fatty acids containing four or more double bonds, when the saponified solution is cooled. They also compared the cooling temperature and the rate of cooling on the enrichment of omega-3 fatty acids of alkali salts of sardine oil fatty acids. The DHA and EPA from sardine oil were concentrated more than 2.3fold with minimum yields of 91 and 87%, respectively. Fatty acid profiles of the prepared concentrates showed that the cooling rate and temperature had little effect on the yield and contents of DHA and EPA. Chen and Ju (2001) utilized a modified low-temperature solvent crystallization process for the enrichment of PUFAs in borage and linseed oil fatty acids. The effects of solvent, operation temperature, and solvent-to-FFA ratio on the concentration of PUFAs were investigated. The best results were achieved when a mixture of 30% acetonitrile and 70% acetone was used as the solvent. With an operating temperature of −80 °C and a solvent-to-FFA ratio of 30 mL g –1, g-linolenic acid (GLA; 18:3n-6) in FFA of saponified borage oil was raised
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490 Separation, extraction and concentration processes from 23.4 to 88.9% with a yield of 62.0%. At a yield of 24.9%, a-linolenic acid (ALA; 18:3n-3) in linseed oil was increased from 55.0 to 85.7%.
17.5 Supercritical-fluid extraction for the concentration and purification of omega-3 fatty acids Supercritical-fluid extraction is a relatively novel technique which has found use in food and pharmaceutical applications (Senanayake, 2000). The properties of a supercritical fluid are used to selectively extract a specific compound or to fractionate mixtures by changing the temperature and pressure without any phase change. The most important property of these fluids is the high solvation capacity in their supercritical region (Shahidi and Senanayake, 2006; Shahidi and Wanasundara, 1998). This method is mild and, because it uses CO2, minimizes autoxidation (Senanayake, 2000). A number of gases are known to have good solvent properties at pressures above their critical values. For food applications, CO2 is the solvent of choice because it is inert, inexpensive, non-flammable, environmentally acceptable, non-toxic, relatively safe, completely recoverable, and readily available and has a moderate critical temperature (31.1 °C) and pressure (1070 psig). The fatty acids were most effectively separated on the basis of chain length; hence the method works best for oils with low levels of long-chain fatty acids. Because this method is derived from separation of compounds based on their molecular weight and not their degree of unsaturation, prior concentration steps, such as urea adduction or low-temperature crystallization, may be necessary in order to purify the omega-3 fatty acids. The use of supercritical fluids for extraction and purification of omega-3 fatty acids from fish and krill oils has been reported (Mishira et al., 1993; Yamagouchi, et al., 1986). Fish oils in the form of free fatty acids and fatty acid esters have been extracted with supercritical CO2 to yield concentrates of EPA and DHA. The use of high pressures and high capital costs might limit the widespread use of this technique in large-scale applications (Shahidi and Senanayake, 2006). Alkio et al. (2000) evaluated the technical and economic feasibility of producing ethyl ester concentrates of DHA and EPA from transesterified tuna oil using supercritical fluid chromatography. A systematic experimental procedure was used to find the optimal values for process parameters and the maximal production rate. DHA ester concentrates up to 95 wt% purity were obtained in one chromatographic step with supercritical fluid chromatography, using CO2 as the mobile phase at 65 °C and 145 bar and octadecylsilane-type reversed-phase silica as the stationary phase. DHA ester, 0.85 g/(kg stationary phase/h) and 0.23 g EPA ester/(kg stationary phase/h) can be simultaneously produced at the respective purities of 90 and 50 wt%. The process for producing 1000 kg of DHA concentrate and 410 kg of EPA concentrate per year requires 160 kg of stationary phase
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Methods of concentration and purification of omega-3 fatty acids 491
and 2.6 t h–1 of CO2 as recycling mobile phase. The operating cost of the supercritical fluid chromatography was US$550 kg–1 for DHA and EPA ethyl ester concentrate. The US Patent 4,675,132 (Stout and Spinelli, 1987) demonstrated that fish oil esters could be fractionated by supercritical fluid extraction to produce an oil with a DHA content of 60–65%. However, the recovery of omega-3 fatty acids was low during this process. Jachmanián et al. (2007) studied the solubility of various ethyl esters derived from hake liver oil in supercritical CO2. A selectivity factor was used to determine optimal conditions to fractionate the ethyl ester mixture. A strong influence of solvent pressure and temperature was observed at 8.63–18.04 MPa and 40–70 °C. The lowest total solubility of the ethyl ester mixture was obtained when using supercritical CO2 at the lowest density (the lowest pressure and the highest temperature value tested). The highest discrimination against longchain PUFAs (e.g. EPA and DHA) was also obtained under these conditions. Conversely, higher solubility and lower selectivity were obtained when solvent density increased. Considering this inverse correlation between selectivity and solubility, a single-step batch-fractionation process was designed to increase the DHA ethyl ester content from an initial value of 17.5% in the starting material to 55% in the final extract. Davarnejad et al. (2008) examined the solubility of fish oil in supercritical CO2 at temperatures of 40, 50, 60, and 70 °C and pressures of 13.6, 20.4, and 27.2 MPa. The fractionated fish oil samples collected were then esterified using methanol with sodium methoxide catalyst. The samples were analyzed by GC to determine the amount of four fatty acid methyl ester (FAME) components extracted, namely, methyl palmitate, methyl oleate, methyl EPA (5,8,11,14,17-eicosapentanoate), and methyl DHA (4,7,10,13,16,19docosahexenoate). The results showed that the highest solubility of the fish oil (0.921 g of oil in 100 g of CO2) was obtained at optimum conditions of 40 °C and 27.2 MPa. The solubility of fish oil in supercritical CO2 was found to be higher at lower temperature and at lower fractionation time. Furthermore, the average yield obtained for the combined total of the four FAME components was 66%, with methyl palmitate having the highest at 30.5% under extraction conditions of 50 °C and 13.6 MPa whereas methyl EPA has the lowest at 3.24%. Létisse et al. (2006), and Létisse and Comeau (2008) evaluated the enrichment of EPA and DHA from sardine heads, a waste product from the fish canning industry, via supercritical fluid extraction. These studies were done on a laboratory scale. Various parameters, such as pressure, temperature, CO2 rate and time were optimized in order to obtain the highest yield of extracted oil with the highest amount of EPA and DHA in the extraction product. In the first approach, the oil yield was measured. Then, a quadratic model with three variables was employed to maximize the EPA and the DHA concentrations. A multicriteria optimization, using the desirability function, was performed to determine the best level for each parameter. Pressure, temperature and CO2 rate were, respectively, set at 300 bar, 75 °C and © Woodhead Publishing Limited, 2010
492 Separation, extraction and concentration processes 2.5 ml min–1 during the 45 min extraction. A yield of 10.36% of extracted oil (compared with the dry material) was achieved with an amount of 10.9% of EPA and 13.0% of DHA (compared with all fatty acids of the extract). These yields were lower than with a traditional solvent extraction. However, the advantages of supercritical fluid extraction were shorter extraction time, prevention of heating, and better organoleptic properties by excluding the use of toxic organic solvents. Catchpole et al. (2000) reported the countercurrent extraction and fractionation of a range of crude fish oils using supercritical CO2 and CO2– ethanol mixtures. Vitamin A palmitate was extracted from model mixtures of cod liver oil and vitamins using pure CO2. The separation factor was low, owing to similar solubilities of the vitamin ester and the oil. Vitamin A was also recovered from cod liver oil ethyl esters–vitamin A mixtures. The separation factor was substantially improved over the non-esterified oil, owing to large differences in the solubilities of the esters and vitamin A in supercritical CO2. Solubilities of fish oils and squalene are reported using CO2–ethanol mixtures at 333 K, ethanol concentrations of 0 to 12% by mass, and pressures of 200–300 bar. Solubilities of all oils and squalene increased exponentially with linear increases in the ethanol concentration. The solubility of polar components increased more rapidly than non-polar components. Pilot-scale removal of fatty acids from Orange Roughy oil and squalene from deep sea shark liver oil was carried out using CO2–ethanol mixture as the solvent. The extent of fatty acid removal from Orange Roughy oil was higher than with pure CO2, whereas the degree of separation of squalene from shark liver oil was lower. However, throughput was substantially increased relative to pure CO2 in both cases. Temelli et al. (1995) optimized the supercritical CO2 extraction temperature and pressure for oil removal from freeze-dried, fall Atlantic mackerel. The effect of extraction conditions on pH and water-binding potential of the protein residue was evaluated. For the temperature range (35–55 °C) and pressure interval (20.7–34.5 MPa), a combination of 34.5 MPa and 35 °C gave the highest oil yield and concentration of omega-3 fatty acids.
17.6 Distillation methods for the concentration and purification of omega-3 fatty acids Fractional distillation is another simple process for separation of mixtures of omega-3 fatty acids under ultralow pressure. This method takes advantage of differences in the boiling point and molecular weight of fatty acids under the conditions of high temperature (180–250 °C) and reduced pressure (0.1 to 1.00 mm Hg). Fatty acid enrichment is achieved by exploiting the differences in vapour pressure through countercurrent contacting of vapour
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and liquid phases in stages using plates or continuously using random or structured packing. Molecular distillation, often also called ‘short path distillation’, may use moderately lower temperatures and short heating intervals. Molecular distillation provides high selectivity, but as a result of its relatively high operating temperatures and reduced pressures, it leads to high operating costs, possible thermal degradation of omega-3 fatty acids and safety concerns. Several advantages and disadvantages of molecular distillation are reported in Table 17.3. Molecular distillation can be used for the concentration of omega-3 fatty acids from fish oils. The content of omega-3 fatty acids in fish oils mainly depends on the fish species itself, their habitat and the season. A typical base material for commercial production of omega-3 fatty acid concentrates is anchovy or sardine oil with a content of approximately 18% EPA and 12% DHA. Tuna oil is also used for the large-scale production of DHA concentrates, as the crude oil contains up to 25% DHA. In general, the concentration of omega-3 fatty acids in the crude oil should not be less than 30% in total. The state-of-the-art method for the concentration of omega-3 fatty acids from fish oil by molecular distillation is by the fractionation of their ethyl esters (Fig. 17.1). The latter can be obtained by interesterification of fish oil with ethyl alcohol. The equipment used in this process consists of two stills (two-stage molecular distillation) with a degasser. The product distillation requires a minimum of three passes through the system. The first, degassing pass, removes any moisture left after interesterification. The residue from this pass is then used as the feed for the second, ‘lights’ removal, pass. The second pass concentrates the DHA and EPA by separating and removing the ‘lights’ from the feed. The second pass removes 20–50% of C10–C18 fatty acid esters. The residue from this pass is then used as the feed for the third, ‘product distillation’, pass. The third pass further concentrates the DHA and EPA (up to 40–80%) by separating the heavier fraction from the feedstock. The heavy fraction (5–10%) consists of longer chain fatty acids, including fatty acids chain length >C24 as residue. Operating pressures and temperatures required are 0.005–0.01 torr and 170–190 °C, respectively. The recoveries can be up to 70% at DHA and EPA concentrations of 55–65 wt%. Table 17.3 Advantages and disadvantages of molecular distillation Advantages
Disadvantages
Stability: the vacuum allows oils to be processed at minimal temperatures, reducing the risk of thermal decomposition and oxidation of PUFAs
Cost: cost is relatively high
Purity: separating the oil components by weight allows contaminants to be reduced far below industry specifications
Natural form: the starting natural triglyceride form is lost in the process
Concentration: weight grouping allows the processor to concentrate specific fatty acids
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494 Separation, extraction and concentration processes Crude fish oil
Refining, bleaching and deodorization
Refined fish oil
Interesterification
Fish oil ethyl esters
Molecular distillation
Purified omega-3 fatty acid esters
Fig. 17.1 Purification of crude fish oil into highly concentrated omega-3 fish oil esters.
The DHA to EPA ratio primarily depends on their content in the base material and on the degree of concentration, as EPA ethyl ester, the more volatile component, also accumulates in the mid-chain-length fatty acid ester fraction. Reprocessing of the latter fraction by molecular distillation increases the recovery of omega-3 fatty acids. Stout et al. (1990) pointed out the practical difficulty of purifying omega-3 fatty acids from menhaden oil via molecular distillation in the natural TAG form. The distillation of menhaden oil, in its natural TAG form, concentrated only EPA from an initial value of 16.0% to 19.5%. Hence, this study did not result in a significant improvement in the concentration of omega-3 fatty acids, as they are more or less uniformly distributed in the TAGs. Furthermore, distillation of triglycerides require high temperatures, risking thermal decomposition of unsaturated fatty acids, and a very low operating pressure, which results in an inefficiently low specific feed rate. However, molecular distillation of its ethyl esters increased the EPA content from 15.9% to 28.4% (Stout et al., 1990). The degree of concentration of DHA was even more remarkable. Whereas DHA doubled from 8.4% to 17.3% in the TAG form, in the alkyl ester form it increased from 9.0% to 43.9% (Stout et al. 1990). Although the greater volatility of the fatty acid esters allows the use of lower temperatures (compared with process temperatures required for FFA fractionation), temperatures are still moderately high, and exposure to distillation conditions over a prolonged period of time can be © Woodhead Publishing Limited, 2010
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detrimental to the omega-3 fatty acid constituents, causing polymerization, hydrolysis, isomerization and thermal degradation. Methyl esters of Atlantic herring oil containing 62% monoethylenic fatty acids were subjected to batch fractional distillation, under vacuum on a pilot-plant scale, to study the feasibility of fractionating fatty acid esters of marine oils of low iodine value into monounsaturated fractions with increased commercial value for industrial chemical uses (Ackman et al., 1973). A total of 64 methyl ester fractions were collected and recoveries of the major saturated and monounsaturated acids were close to 100%, and some fractions contained over 90% of the desired 22:1 long chain monounsaturated acids. The short-chain polyunsaturated acids were recovered in high yields, but the long-chain highly unsaturated acids were recovered in yields of 60% or less owing to oxidative and thermal decomposition in the particular apparatus employed. If small amounts of unsaturated acids are acceptable, fractional distillation of low iodine value marine oils could inexpensively supply fractions with high concentrations of methyl esters of longer chain (C20 and C22) monounsaturated and shorter chain (C14) saturated acid or (C16) saturated–monounsaturated acid mixture. Liang and Hwang (2000) employed short-path distillation to fractionate EPA and DHA ethyl esters from squid visceral oil. The elimination temperatures of squid visceral oil ethyl esters (SVOEE) ranged from 50 to 140 °C, increasing with the carbon number of ethyl esters. The elimination temperature of cholesterol was higher than those of SVOEE. When SVOEE originally containing 9.0% EPA, 14.7% DHA, and 11.21 mg g–1 of cholesterol was distilled from 50 to 150 °C, the 130 °C distillate contained 15.5% EPA and 34.7% DHA with 0.99 mg g–1 of cholesterol, and the yield was 21.8%. The 150 °C distillate had 43.1% DHA with 4.96 mg g–1 of cholesterol. Furthermore, the distillates collected from 110 to 150 °C contained 24.4 to 50.2% of EPA plus DHA, and their total yield was 58.3%.
17.7 Enzymatic methods for the concentration and purification of omega-3 fatty acids For the concentration of omega-3 fatty acids on a large scale, each of the above physical and chemical separation methods have some drawbacks either in terms of low yield, a requirement for large volumes of solvent or sophisticated equipment, a risk of structural changes in the fatty acid products, or high operational costs (Senanayake, 2000). Lipases work under mild conditions of temperature and pH (Gandhi, 1997), a factor which favors their potential use for the enrichment of omega-3 fatty acids in oils. Lipases are enzymes that catalyze the hydrolysis, esterification, interesterification, acidolysis and alcoholysis reactions (Senanayake and Shahidi, 2000b; Shahidi and Senanayake, 2006; Shimada et al., 2006; Weete et al., 2008). The common
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496 Separation, extraction and concentration processes feature among lipases is that they are activated by an interface. Lipases have been used for many years to modify the structure and composition of foods. Lipases which act on neutral lipids generally hydrolyze the esters of PUFAs at a slower rate than those of more saturated fatty acids (Villeneuve and Foglia, 1997). Use has been made of this relative substrate specificity to increase the concentration of omega-3 PUFAs in seal blubber and menhaden oils by subjecting them to hydrolysis by a number of microbial lipases (Wanasundara and Shahidi, 1998). Concentration of omega-3 fatty acids by enzyme-assisted reactions involves benign reaction conditions and provides an alternative to the traditional concentration methods such as distillation and chromatographic separation (Shahidi and Senanayake, 2006). Furthermore, concentration via enzymatic means may also produce omega-3 fatty acids in the acylglycerol form, which is nutritionally preferred. Studies over the past two decades have used microbial lipases to produce purified omega-3 fatty acids via hydrolysis, esterification or transesterification of marine oils. Tanaka et al. (1992) used six types of microbial lipases (lipases derived from Aspergillus niger, Candida cylindracea, Chromobacterium viscosum, Rhizopus javanicus, Rhizopus delemer and Pseudomonas sp.) to hydrolyze tuna oil and found that the lipase from Candida cylindracea was the most effective one in increasing the DHA content in the non-hydrolyzed fraction of the oil. The DHA content in the non-hydrolyzed fraction was increased three-fold compared with the original oil; however, other lipases were not very effective. Various microbial lipases were evaluated for the enrichment of omega-3 fatty acids from cod liver and sardine oils by selective hydrolysis (Hoshino et al., 1990). The best hydrolysis results were obtained for the lipases from Candida cylindracea and Aspergillus niger, but none of the lipases were able to increase the EPA content of the oil considerably. However, over 50% of the total omega-3 fatty acids were produced when these two lipases were used. A Japanese patent (Noguchi and Hibino, 1984) describes a method based on the discrimination of lipases on DHA and EPA for preparation of omega-3 fatty acids. Ethyl esters from sardine and mackerel were hydrolyzed with several lipases derived from Candida cylindracea, Mucor miehei and Aspergillus niger. Selective hydrolysis afforded ethyl ester concentrates of up to 17% DHA and 25% EPA after separation of the hydrolyzed fatty acids. Transesterification of various fish oil TAGs with a stoichiometric amount of ethanol catalyzed by immobilized Rhizomucor miehei lipase under anhydrous solvent-free conditions resulted in a good separation of EPA and DHA (Haraldsson and Kristinsson, 1998). When FFAs from the various fish oils were directly esterified with ethanol under similar conditions, greatly improved results were obtained. When tuna oil comprising 6% EPA and 23% DHA was transesterified with ethanol, 65% conversion into ethyl esters was obtained. The residual glyceride mixture contained 49% DHA and 6% EPA, with 90% DHA recovery into the glyceride mixture and 60% EPA recovery into the ethyl ester product. When the corresponding tuna oil FFAs were directly esterified with ethanol, 68% conversion was obtained.
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The residual FFAs comprised 74% DHA and only 3% EPA. The recovery of both DHA into the residual FFA fraction and EPA into the ethyl ester product was reported to be very high at 83 and 87%, respectively (Haraldsson and Kristinsson, 1998). Schmitt-Rozieres et al. (2000) studied the enrichment of PUFAs from sardine cannery effluents via enzymatic selective esterification. The sardine canning industry produces vast quantities of effluents that need expensive reprocessing. Their oily component contains EPA and DHA up to 10% each. The author’s goal was to develop a process allowing the recovery of these fatty acids. After the removal of solid particles, proteins, and peptides from the crude effluent, the resultant oil was hydrolyzed and EPA and DHA were enriched by selective enzymatic esterification. Using Lipozyme™, DHA was enriched up to 80%, but no enrichment was observed for EPA. By immobilizing Candida rugosa lipase on Amberlite IRC50 cation-exchange resin, a 30% EPA enrichment was obtained. Shimada et al. (2001) also attempted to purify PUFAs by taking advantage of the enzyme-catalyzed reactions. When FFAs originating from PUFA-containing oil were selectively esterified with lauryl alcohol (LauOH) using a lipase acting on the desired PUFAs very weakly, the PUFA was efficiently enriched in the FFA fraction. In addition, when selective alcoholysis of ethyl esters originating from PUFA-containing oil with LauOH was carried out, the PUFA ethyl ester (EtPUFAs) was enriched to a desired purity in the unreacted ethyl ester fraction. These reaction mixtures contain LauOH, PUFA (EtPUFAs), and lauryl esters, and their molecular weights are different from one another. Hence, PUFA or EtPUFAs can be easily separated by conventional distillation. Selective esterification increased the purity of DHA, GLA, and arachidonic acid (ARA; 20:4n-6) to 91, 98, and 96 wt%, respectively. Selective alcoholysis was also effective for increasing the purity of ethyl docosahexaenoate to 90 wt%. Kojima et al. (2006) studied the enzymatic acidolysis and acylglycerol synthesis using PUFAs with lipases from Pseudomonas fluorescens HU380 (HU-lipase), P. fluorescens AK102 (AK-lipase), and Candida rugosa (CR-lipase). The acidolysis of triolein with EPA or DHA in n-hexane was evaluated with lipases immobilized on Celite 545. HU-lipase showed the highest incorporation rate at a low temperature (10 °C) with either EPA or DHA as the acyl donor, and the rate decreased with increasing reaction temperature. At 45 °C, the rates for EPA and DHA were 7.1 and 0.5 relative to those at 10 °C, respectively. The EPA incorporation rate was even higher at a low temperature (10 °C), and the DHA incorporation rate increased with decreasing temperature. Although AK-lipase showed the reverse tendency for incorporation rate, the DHA incorporation rate increased with increasing reaction temperature with both PUFAs. HU-lipase reacted well with PUFAs such as DHA, EPA, ARA, mead acid (MA), and dihomo-g-linolenic acid (DGLA) on acidolysis and glyceride synthesis. The reactivities of AK-lipase toward these PUFAs except for DGLA, i.e. MA, ARA, EPA, and DHA, were low for both reactions.
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498 Separation, extraction and concentration processes Halldorsson et al. (2003) investigated the use of lipases as catalysts for separating EPA and DHA in fish oil by kinetic resolution based on their fatty acid selectivity. Esterification of FFA from various types of fish oils with glycerol by immobilized Rhizomucor miehei lipase under water-deficient, solvent-free conditions resulted in a highly efficient separation of EPA and DHA. Reactions were conducted at 40 °C with a 10% dosage of the lipase preparation under vacuum to remove the coproduced water, thus rapidly shifting the reaction toward the products. The bulk of the fatty acids, together with EPA, were converted into acylglycerols, whereas DHA remained in the residual FFA. When FFA from tuna oil comprising 5% EPA and 25% DHA were esterified with glycerol, 90% conversion into acylglycerols was obtained. The residual FFA contained 78% DHA and only 3% EPA, with 79% DHA recovery. EPA recovery in the acylglycerol fraction was 91%. The type of fish oil and extent of conversion were highly important parameters in controlling the degree of concentration of EPA and DHA.
17.8 Integrated methods for the concentration and purification of omega-3 fatty acids Most of the separation methods described above, when used alone, can only concentrate and purify omega-3 fatty acids to a limited extent. Therefore, two or more procedures are often required to produce highly purified omega-3 fatty acids. A simple and inexpensive method involving saponification of wet biomass, followed by transmethylation, winterization and urea adduction in a sequential manner has been recently developed for concentration of DHA from Crypthecodinium cohnii biomass (Mendes et al., 2007) (Fig. 17.2). The algal biomass grown in shake flasks is concentrated by centrifugation and the wet concentrate is kept frozen until needed. Fatty acids are extracted by direct saponification of wet biomass with KOH-ethanol. Before extracting unsaponifiable matter with hexane, water is added to shift the equilibrium distribution of unsaponifiable matter to the hexane phase. The hydroalcoholic phase, containing the soaps, is acidified with HCl and free fatty acids are extracted with hexane. The organic phase, containing FFAs, is dried with anhydrous sodium sulfate and the solvent is evaporated to recover FFAs. The FFAs are then methylated using transmethylation reagent and sulfuric acid. The mixture containing methyl esters are then winterized at –18 °C. Subsequently, the liquid fraction is separated from the winterized crystals. The liquid fraction is added to a urea-methanol saturated solution and the urea-fatty acids adducts are crystallized; the crystals are separated by centrifugation and the filtrate (non-urea complexing fraction), which contains DHA, is finally recovered. The highest DHA recovery (49.9%) was obtained at 24 °C at a urea/fatty acid ratio of 4.0, corresponding to 89.4% of DHA of the total fatty acids (Mendes et al., 2007).
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Methods of concentration and purification of omega-3 fatty acids 499 Microalgal wet biomass Saponification Free fatty acids Transmethylation Fatty acid methyl esters (FAMEs) Winterization DHA methyl esters Urea adduction Highly purified DHA methyl esters
Fig. 17.2 Scheme for the DHA concentration and purification of microalgal biomass from Crypthecodinium cohnii.
Medina et al. (1995) described a two-step process, urea adduction followed by preparative high-performance liquid chromatography, for concentration of omega-3 fatty acids from the marine microalga, Isochrysis galbana. By the urea adduction method, a mixture that contained 94% (w/w) stearidonic (SA), EPA, and DHA acids (4:1 urea/fatty acid ratio and 4 °C crystallization final temperature) was obtained from cod liver oil fatty acids. Further purification of SA, EPA, and DHA was achieved with reverse-phase C18 columns. These isolations were scaled up to a semi-preparative column. A PUFA concentrate was isolated from I. galbana with methanol/water (80:20, w/w) or ethanol/ water (70:30, w/w). With methanol/water, a 96% EPA fraction with 100% yield was obtained, as well as a 94% pure DHA fraction with a 94% yield. With ethanol/water as the mobile phase, EPA and DHA fractions obtained were 92% pure with yields of 84 and 88%, respectively. Guil-Guerrero and Belarbi (2001) also used an integrated method to purify DHA and EPA from cod liver oil. The process consisted of four main steps: (i) saponification of the oil, (ii) use of urea inclusion adducts method to obtain PUFAs, (iii) PUFA methylation, and (iv) silica gel column chromatography of the methylated PUFAs. Silica gel chromatography yielded highly pure DHA in the process (100% purity, 64% yield). For EPA, the recovery in the combined process was 29.6%, and the final purity was 90.6%, owing to the simultaneous elution of other PUFA esters. The recovery in the urea adduction method was strongly enhanced by application of orbital agitation during the crystallization process, in which EPA yield increased from 60–70% without agitation to 90–97% at 800 rpm; stearidonic acid (18:4n−3) yield ranged from 60–75% without agitation to 87–95% at 800 rpm, and DHA yield varied from 53–73% without agitation to 85–99% at 800 rpm. Chakraborty and Paulraj (2008) used enzymatic hydrolysis followed © Woodhead Publishing Limited, 2010
500 Separation, extraction and concentration processes by urea adduction to purify EPA and ARA from sardine oil. The enzyme used for the hydrolysis of sardine oil was the lipase derived from Bacillus licheniformis MTCC 6824. The enzyme exhibited more hydrolytic resistance toward the ester bonds of EPA and ARA than those of other fatty acids and was proved to be effective for increasing the concentration of EPA and ARA from sardine oil. Utilizing this fatty acid specificity, EPA and ARA from sardine oil were enriched by lipase-mediated hydrolysis followed by urea fractionation at 4 °C. The purified lipase produced the highest degree of hydrolysis for SFAs and MUFAs (81.5 and 72.3%, respectively, from their initial content in sardine oil) after 9 h. The profile of conversion by lipase catalysis showed a steady increase up to 6 h and thereafter plateaued. Lipasecatalyzed hydrolysis of sardine oil followed by urea adduction with methanol provided FFAs containing 55.4% EPA and 5.8% ARA, respectively, after complexation of saturated and less unsaturated fatty acids. The combination of enzymatic hydrolysis and urea adduction proved to be a promising method to obtain highly concentrated EPA and ARA from sardine oil. Yuzo et al. (2006) employed enzymatic hydrolysis, using lipase from Pseudomonas fluorescens strain HU380, and urea adduction to concentrate EPA and DHA from refined cod oil. The starting oil had 12.2% EPA and 6.9% DHA. Lipase-catalyzed hydrolysis followed by urea adduction provided FFAs with 43.1% EPA and 7% DHA. The resulting yield of concentrated total fatty acids comprised 2.6% of the fatty acids from the cod oil. Thus, EPA was particularly concentrated in the fatty acids derived from refined cod oil as a result of enzymatic hydrolysis followed by urea adduction. On the other hand, hydrolysis of cuttlefish oil with AK-lipase (lipase from Pseudomonas fluorescens strain AK102), followed by urea adduction increased the EPA content from 14.2 to 16.8%, and DHA content from 16.3 to 44.6%. The yield of purified total fatty acids by urea concentrate was 9.4% of the fatty acids from the cuttlefish oil. Thus, DHA was particularly concentrated in this study. A US patent (Zaks and Gross, 1999) disclosed an enzymatic process for preparing an oil-based product significantly enriched in omega-3 fatty acids. The process is a two-step procedure involving lipase-catalyzed transesterification of TAGs followed by low-temperature crystallization. The process yields a mixture of highly pure monoglycerides, at least 60% of which contain omega-3 fatty acids. The process can also be used to produce omega-3 enriched TAG products. Enzymatic methods of enrichment of omega-3 fatty acids have also been investigated under supercritical conditions. Lin et al. (2006) examined the enrichment of omega-3 fatty acids in TAGs of menhaden oil by lipasecatalyzed trans-esterification under supercritical CO2. Before the reaction, menhaden oil was treated by urea adduction to produce 80.1% of omega-3 PUFAs containing 29.4% EPA and 41.8% DHA. Using the sn-1,3-specific lipase from Mucor miehei, the effect of various operating parameters on the reaction was studied, including co-solvent concentration, reaction temperature,
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Methods of concentration and purification of omega-3 fatty acids 501
time, pressure, and substrate ratio (free omega-3 fatty acids: TAGs). Water and ethanol were examined as co-solvents, and both fluids exhibited a maximum for omega-3 fatty acid content in TAGs as a function of concentration. Pressure up to 103.4 bar had a significant positive effect on the conversion of omega-3 PUFA onto TAGs. By further increasing pressure, the conversion rate decreased owing to the transformation of the spatial structure of lipase that leads to deactivation. The enzyme exhibited a good performance and stability in the region of 323 K. The optimal substrate molar ratio of TAG to omega-3 PUFA is about 1:4 taking into consideration the TAG inhibition. The conversion in supercritical CO2 appeared 40% higher than in n-hexane at ambient pressure after 5h. In another study, Lin and Chen (2008) used esterification of free omega-3 fatty acids with enriched omega-6 TAGs to produce a desired structured lipid with an omega-3/omega-6 ratio of 4 by using lipases under supercritical CO2. Comparing four different types of lipases, sn-1,3-specific lipase from Mucor miehei had the highest degree of incorporation under 10 wt% loading amount of total substrates. The optimal operating parameters under 10.2 MPa and 323.15 K supercritical CO2 could attain the desired omega-3/ omega-6 ratio in 6 h. Because of the negative effect on the enzyme activity by the enriched omega-6 TAGs, the optimal substrate ratio of the enriched omega-6 TAGs and the omega-3 fatty acids was chosen as 1/4. To enhance the solubility of omega-3 fatty acids in supercritical CO2, ethanol was applied as a co-solvent and reached an optimal input at 10% of total substrates. The enzyme maintained 81% of its initial activity because of moisture removal from the surface of the enzyme after seven cyclic pressurization/ depressurizations.
17.9 Conclusions Production of purified omega-3 fatty acids from natural source materials may be achieved via a number of techniques, namely urea adduction, chromatography, low-temperature crystallization, supercritical fluid extraction, enzymatic splitting, molecular distillation, as well as a combination of any of the above methods. Purified omega-3 fatty acid products thus produced may be in the form of free fatty acids, alkyl esters or acylglycerols. Each method had its own benefits and drawbacks. Of the methods described above, molecular distillation is highly energy consuming and results in a significant distraction of labile highly unsaturated fatty acids. Other methods, which require the use of organic solvents, include low-temperature crystallization, chromatography and supercritical CO2 extraction, among others. These processes have a number of drawbacks. For example, low-temperature crystallization of acylglycerols typically results in only a small omega-3 enrichment of the product, and chromatography and supercritical fluid extraction are expensive and difficult to scale up. Owing to the potential benefits of having the concentrates in the © Woodhead Publishing Limited, 2010
502 Separation, extraction and concentration processes acylglycerol form, enzymatic procedures have become popular. However, there are several limitations to the enzymatic processes as well. These include the necessity for a complex separation of the product from free saturated fatty acids, use of complex multi-enzymatic systems and the low efficiency that results in an insufficient degree of enrichment. Urea adduction has been very successful in enriching omega-3 fatty acids on an industrial scale and the products thus formed are in the free fatty acid or simple ester forms. Efficient and cost-effective methods of enriching the level of omega-3 fatty acids will continue to be needed in order to reduce the cost and to meet the future demand for highly purified omega-3 products.
17.10 References Ackman R G, Ke P J and Jangaard P M (1973), ‘Fractional vacuum distillation of herring oil methyl esters’, J Am Oil Chem Soc, 50, 1–8. Alkio M, Gonzalez C, Jäntti M and Aaltonen O (2000), ‘Purification of polyunsaturated fatty acid esters from tuna oil with supercritical fluid chromatography’, J Am Oil Chem Soc, 77, 315–321. Bousquet O and Goffic F L (1995), ‘Countercurrent chromatographic separation of polyunsaturated fatty acids’, J Chromatogr, 704, 211–216. Catchpole O J, Grey J B and Noermark K A (2000), ‘Fractionation of fish oils using supercritical CO2 and CO2+ethanol mixtures’, J Supercrit Fluids, 19, 25–37. Chakraborty K and Paulraj R (2008), ‘Enrichment of eicosapentaenoic acid from sardine oil with D5-olefinic bond specific lipase from Bacillus licheniformis MTCC 6824’, J Agric Food Chem, 56, 1428–1433. Chen T and Ju Y (2001), ‘Polyunsaturated fatty acid concentrates from borage and linseed oil fatty acids’, J Am Oil Chem Soc, 78, 485–488. Davarnejad R, Kassim K M, Zainal A and Sata S A (2008), ‘Extraction of fish oil by fractionation through supercritical carbon dioxide’, J Chem Eng Data, 53, 2128– 2132. Du Q, Shu A and Ito Y (1996), ‘Purification of fish oil ethyl esters by high-speed countercurrent chromatography using non-aqueous solvent system’, J Liq Chromatogr Relat Technol, 19, 1451–1457. Gandhi N (1997), ‘Applications of lipase’, J Am Oil Chem Soc, 74, 621–634. Guil-Guerrero J L and Belarbi E (2001), ‘Purification process for cod liver oil polyunsaturated fatty acids’, J Am Oil Chem Soc, 78, 477–484. Haagsma N, Gent C M, Luten J B, Jong R W and Doorn E (1982), ‘Preparation of an w3 fatty acid concentrate from cod liver oil’, J Am Oil Chem Soc, 59, 117–118. Halldorsson A, Kristinsson B, Glynn C and Haraldsson G G (2003), ‘Separation of EPA and DHA in fish oil by lipase-catalyzed esterification with glycerol’, J Am Oil Chem Soc, 80, 915–921. Han D S, Ahn H B and Shin H K (1987), ‘Separation of EPA and DHA from fish oil by solubility differences of fatty acid salts in ethanol’, Korean J Food Sci Technol, 19, 430–434. Haraldsson G G and Kristinsson B (1998), ‘Separation of eicosapentaenoic acid and docosahexaenoic acid in fish oil by kinetic resolution using lipase’, J Am Oil Chem Soc, 75, 1551–1556. Hayashi K and Kishimura H (1993), ‘Preparation of n-3 PUFA ethyl ester concentrates from fish oil by column chromatography on silicic acid’, Nippon Suisan Gakkaishi, 59, 1429.
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Methods of concentration and purification of omega-3 fatty acids 503 Hayes D G (2006), ‘Effect of temperature programming on the performance of urea inclusion compound-based free fatty acid fractionation’, J Am Oil Chem Soc, 83, 253–259. Hayes D G, Alstine J M V and Setterwall F (2000), ‘Urea-based fractionation of seed oil samples containing fatty acids and acylglycerols of polyunsaturated and hydroxy fatty acids’, J Am Oil Chem Soc, 77, 207–213. Hoshino T, Yamane T and Shimuzu S (1990), ‘Selective hydrolysis of fish oil by lipase to concentrate w3-polyunsaturated fatty acids’, Agric Biol Chem, 54, 1459–1467. Hwang L S and Liang J (2001), ‘Fractionation of urea-pretreated squid visceral oil ethyl esters’, J Am Oil Chem Soc, 78, 473–476. Jachmanián I, Margenat L, Torres A I and Grompone M A (2007), ‘Selectivity of supercritical CO2 in the fractionation of hake liver oil ethyl esters’, J Am Oil Chem Soc, 84, 597–601. Kojima Y, Sakuradani E and Shimizu S (2006), ‘Acidolysis and glyceride synthesis reactions using fatty acids with two Pseudomonas lipases having different substrate specificities’, J Biosci Bioeng, 102, 179–183. Kris-Etherton P, Eckel R H, Howard B V, St Jeor S and Bazzarre T L (2001), ‘AHA Science Advisory: Lyon diet heart study. Benefits of a Mediterranean-style, national cholesterol education program/American Heart Association Step I dietary pattern on cardiovascular disease’, Circulation, 103, 1823–1825. Lands, W E M (2003), ‘Diets could prevent many diseases’. Lipids, 38, 317–321. Létisse M and Comeau L (2008), ‘Enrichment of eicosapentaenoic acid and docosahexaenoic acid from sardine by-products by supercritical fluid fractionation’, J Sep Sci, 31, 1374–1380. Létisse M, Rozieres M, Hiol A, Sergent M and Comeau L (2006), ‘Enrichment of EPA and DHA from sardine by supercritical fluid extraction without organic modifier I. Optimization of extraction conditions’, J Supercrit Fluid 38, 27–36. Liang J and Hwang L S (2000), ‘Fractionation of squid visceral oil ethyl esters by shortpath distillation’, A Am Oil Chem Soc, 77, 773–777. Lin, T and Chen, S (2008), ‘Enrichment of n-3 polyunsaturated fatty acids into acylglycerols of borage oil via lipase-catalyzed reactions under supercritical conditions’, Chem Engi J, 141, 318–326. Lin T, Chen S and Chang A (2006) ‘Enrichment of n-3 PUFA contents on triglycerides of fish oil by lipase-catalyzed transesterification under supercritical conditions’, Biochem Eng J, 29 27–34. Mansour M P (2005), ‘Reversed-phase high-performance liquid chromatography purification of methyl esters of C(16)–C(28) polyunsaturated fatty acids in microalgae, including octacosaoctaenoic acid [28:8(n-3)]’, J Chromatogr A, 1097, 54–58. Medina R, Giménez A G, Camacho F G, Pérez J A S, Grima E M and Gómez A C (1995), ‘Concentration and purification of stearidonic, eicosapentaenoic, and docosahexaenoic acids from cod liver oil and the marine microalga Isochrysis galbana’, J Am Oil Chem Soc, 72, 575–583. Mendes A, da Silva T L and Reis A (2007), ‘DHA concentration and purification from the marine heterotrophic microalga Crysthecodinium cohnii CCMP 316 by winterization and urea complexation’, Food Technol Biotechnol, 45, 38–44. Merck Index (1983), An encyclopedia of chemical, drugs and biologicals. 10th edn., Rahway, NJ, Merck and Co., Inc. Mishira V K, Temelli F and Ooraikul B (1993), ‘Extraction and purification of omega-3 fatty acids with an emphasis on supercritical fluid extraction’, Food Res Int, 26, 217–226. Murayama W, Kosuge Y, Nakaya N, Nunogaki Y, Nunogaki K, Cazes J and Nunogaki H (1988), ‘Preparative separation of unsaturated fatty acid esters by centrifugal partition chromatography’, J Liq Chromatogr, 19, 283–300. Noguchi Y and Hibino H (1984), ‘Highly unsaturated fatty acid lower ester concentration and separation process’, Japanese Patent, 59–14793. © Woodhead Publishing Limited, 2010
504 Separation, extraction and concentration processes Perretti G, Motori A, Bravi E, Favati F, Montanari L and Fantozzi P. (2007), ‘Supercritical carbon dioxide fractionation of fish oil fatty acid ethyl esters’, J Supercrit Fluid, 40, 349–353. Schmitt-Rozieres M, Deyris V and Comeau L (2000), ‘Enrichment of polyunsaturated fatty acids from sardine cannery effluents by enzymatic selective esterification’, J Am Oil Chem Soc, 77, 329–332. Senanayake S P J N (2000), ‘Enzyme-assisted synthesis of structured lipids containing long-chain omega-3 and omega-6 polyunsaturated fatty acids’, PhD thesis, St. John’s, NF, Canada, Memorial University of Newfoundland. Senanayake S P J N and Shahidi F (2000a), ‘Concentration of docosahexaenoic acid (DHA) from algal oil via urea complexation’, J Food Lipids, 7, 51–61. Senanayake S P J N and Shahidi F (2000b), ‘Structured lipids containing long-chain omega-3 polyunsaturated fatty acids’, in Shahidi F, Seafood in health and nutrition. Transformation in fisheries and aquaculture: global perspectives, St. John’s, NF, Canada, ScienceTech Publishing Co, 29–44. Shahidi F and Senanayake, S P J N (2006). Nutraceuticals and specialty lipids, in Shahidi F, Nutraceuticals and specialty lipids and their co-products, Boca Raton, FL, CRC Press, 1–25. Shahidi F and Wanasundara U N (1998), ‘Omega-3 fatty acid concentrates: nutritional aspects and production technologies’, Trends Food Sci Technol, 9, 230–240. Shimada Y, Nagao T and Watanabe Y (2006), ‘Application of multistep reactions with lipases to the oil and fat industry’, in Shahidi F, Nutraceuticals and specialty lipids and their co-products, Boca Raton, FL, CRC Press, 365–386. Shimada Y, Sugihara A and Tominaga Y (2001), ‘Enzymatic purification of polyunsaturated fatty acids’, J Biosci Bioeng, 91, 529–538. Stout V F, Niisson W B, Krzynowek J and Schlenk H (1990), ‘Fractionation of fish oil and their fatty acids’, In Stansby M E, Fish oils in nutrition, New York, Van Nostrand Reinhold, 73–119. Stout V F and Spinelli J (1987), ‘Polyunsaturated fatty acids from fish oils’, US Patent 4,675,132. Tanaka Y, Hirano J and Funada T (1992), ‘Concentration of docosahexaenoic acid in glyceride by hydrolysis of fish oils with Candida cylindracea lipase’, J Am Oil Chem Soc, 69, 1210–1214. Teshima S, Kanazawa A and Tokiwa S (1978), ‘Separation of polyunsaturated fatty acids by column chromatography on silver nitrate-impregnated silica gel’, Bull J Soc Sci Fish, 44, 927. Temelli F, Leblance E and Long F (1995), ‘Supercritical CO2 extraction of oil from Atlantic mackerel (Scomber scombrus) and protein functionality’, J Food Sci, 60, 703–706. United States Food and Drug Administration (2004). ‘FDA announces qualified health claims for omega-3 fatty acids’. Press release. http://www.fda.gov/bbs/topics/news/2004/ NEW01115.html. Retrieved 09-10-2009. Villeneuve P and Foglia T A (1997), ‘Lipase specificities: Potential application in lipid bioconversions’, Inform, 8, 640–650. Wanasundara U N (1996), ‘Marine oils: stabilization, structural characterization and omega-3 fatty acid concentration’, PhD thesis, St. John’s, NF, Canada, Memorial University of Newfoundland. Wanasundara U N and Fedec P (2002), ‘Centrifugal partition chromatography (CPC): emerging separation and purification technique for lipids and related compounds’, Inform, 13, 726–730. Wanasundara U N and Shahidi F (1998), ‘Lipase-assisted concentration of n-3 polyunsaturated fatty acids in acylglycerols from marine oils’, J Am Oil Chem Soc, 75, 945–951. Wanasundara U N and Shahidi F (1999), ‘Concentration of omega-3 polyunsaturated fatty
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Methods of concentration and purification of omega-3 fatty acids 505 acids of seal blubber oil by urea complexation: Optimization of reaction conditions’, Food Chem, 65, 41–49. Weete J D, Lai O and Akoh C C (2008), ‘Microbial lipases’ in Akoh C C and Min D B, Food lipids: chemistry, nutrition and biotechnology, Boca Raton, FL, CRC Press, 767–806. Yamagouchi K, Murakami W, Nakano H, Konosu S, Kokura T, Yamamoto H, Kosaka M and Hata K (1986), ‘Supercritical carbon dioxide extraction of oils from Antarctic krill’, J Agric Food Chem, 34, 904–907. Yokoyama M, Origasa H, Matsuzaki M, Matsuzawa Y, Saito Y, Ishikawa Y, Oikawa S, Sasaki J, Hishida H, Itakura H, Kita T, Kitabatake A, Nakaya N, Sakata T, Shimada K and Shirato K (2007), ‘Effects of eicosapentaenoic acid on major coronary events in hypercholesterolaemic patients (JELIS): a randomised open-label, blinded endpoint analysis’, Lancet, 369, 1090–1098. Yuzo K, Eiji S and Sakayu S (2006), ‘Different specificity of two types of Pseudomonas lipases for C20 fatty acids with a D5 unsaturated double bond and their application for selective concentration of fatty acids’, J Biosci Bioeng, 101, 496–500. Zaks A and Gross A T (1999), ‘Enzymatic production of monoglycerides containing omega-3 unsaturated fatty acids’, US Patent 5935828. Zuta C P, Simpson B K, Chan H M and Phillips L (2003), ‘Concentrating PUFA from mackerel processing waste’, J Am Oil Chem Soc, 80, 933–936.
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506 Separation, extraction and concentration processes
18 Extraction of natural antioxidants from plant foods E. Conde, A. Moure, H. Domínguez and J. C. Parajó, University of Vigo, Spain
Abstract: An overview of the most studied vegetal sources of compounds with antioxidant properties and their biological action is firstly presented. The conventional and alternative extraction processes are discussed, special emphasis is given to those using biorenewable, environmentally friendly solvents. Selected examples of integrated processes for extraction, concentration and purification of extracts are included. Key words: antioxidants, plants, extraction, purification.
18.1 Introduction The utilization of phytochemicals in the food and cosmetic industries has attracted public and scientific interest, because of their perceived efficiency, low cost and lack of toxicity. The number of recent publications and patents on antioxidants has increased considerably. The demand for natural antioxidants, which are presumed to be safer, has risen owing to concerns about the long-term safety and negative consumer perception of some synthetic antioxidants. Many natural compounds show antioxidant activity and may act as flavorings, colorants, preservatives and reinforcers of endogenous antioxidant systems (Astley, 2003; Valenzuela et al., 2003; Halliwell, 2002; Halliwell and Gutteridge, 1999). There is epidemiological and clinical evidence relating antioxidant-rich diets with a decreased risk of degenerative diseases, reduced morbidity and mortality. However, the safety limits of many natural antioxidants (which can consist of a mixture of several active substances) are unknown, and their lack of toxicity should be confirmed (Galati and
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Extraction of natural antioxidants from plant foods 507 O´Brien, 2004; Pokorny, 2007). The biological activity and bioavailability of natural antioxidants are the subject of much research. Edible vegetals and agroindustrial residues are considered to be abundant and promising sources of natural antioxidants. The extraction method, of crucial importance for both technical and economic reasons, ideally should be non-destructive, time efficient and suitable for producing high quantities of extracts, which should be processed by selective techniques to yield concentrates of enhanced antioxidant capacity. This chapter presents an overview of the most studied vegetal sources of antioxidants and extraction processes (in particular those using biorenewable, environmentally friendly solvents). Selected examples of integrated processes for extraction, concentration and purification of extracts are also included.
18.2 Antioxidant activity in food systems According to a classic definition, an antioxidant is any substance that, when present at low concentrations compared with those of an oxidizable substrate (such as lipids, proteins, DNA or carbohydrates), significantly delays or prevents oxidation of that substrate (Halliwell, 2002; Halliwell and Gutteridge, 1999). In food systems, the term antioxidant is used to designate the inhibitors of lipid peroxidation, whereas in biological systems it usually refers to protection of lipids, proteins and DNA against oxidative damage by processes or reactions involving reactive oxygen and nitrogen species (ROS and RNS, respectively). Antioxidant activity and antioxidant capacity, although often used interchangeably, have different meanings. Activity refers to the rate constant of the reaction between the antioxidant and the oxidant species, and capacity refers to the amount (in moles) of a given free radical scavenged by a sample. The basic mechanism of lipid peroxidation (Frankel, 2005; Laguerre et al., 2007; Pokorny, 1991) is presented in Fig. 18.1. Lipid peroxidation is induced by oxygen in the presence of initiators such as heat, free radicals, light, photosensitizing pigments and metal ions. Once these free radicals are formed, lipid peroxidation progresses at a high rate by a radical chain reaction, and oxidation ends by producing secondary non-radical compounds. In this scheme, antioxidants can act at different levels: ∑ ∑ ∑ ∑ ∑
scavenging species responsible for oxidation initiation or preventive antioxidants; interruptors of the propagation of the autoxidation chain reaction or chain-breaking antioxidants; singlet oxygen quenchers; synergists or compounds increasing the activity of chain-breaking antioxidants in a mixture; reducing agents;
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508 Separation, extraction and concentration processes Initiation Formation of radicals: peroxyl (RO2•), alkoxy (RO•) or alkyl (R•) Propagation R• + O2 Æ RO2• RO2• + RH Æ ROOH + R• RO• + RH Æ ROH + R•
ROOH Æ RO• + •OH 2 ROOH Æ RO2• + RO• + H2O Termination 2 R• R• + RO2• Æ Stable products 2 RO2•
Fig. 18.1 Mechanism of lipid oxidation.
∑ ∑
metal chelators, which stabilize metal pro-oxidants (iron or copper cations); inhibitors of specific oxidative enzymes (especially lipoxygenases).
Antioxidants often act by several mixed and co-operative mechanisms. An antioxidant can behave as a pro-oxidant depending on the structure, chemical environment and operational conditions: for example, typical antioxidants such as flavonoids, a-tocopherol or ascorbic acid can act as pro-oxidants in the presence of transition metal ions, whereas carotenoids show pro-oxidant activity at high oxygen pressure (Galati and O’Brien, 2004). Table 18.1 lists representative in vitro tests employed in the assessment of antioxidant properties. The need for developing simple and reliable in vitro analytical antioxidant tests is widely acknowledged, and the convenience of a common standard has been claimed. A mixture of caffeic acid, catechin and epigallocatechin3-gallate, hesperetin, and morin was proposed for this purpose (Luthria and Vinyard, 2008). The problems associated with the determination of antioxidant activity were summarized in several critical reviews (Andrade et al., 2008; Becker et al., 2004; Decker et al., 2005; Fernandez-Panchón et al., 2008; Frankel and Finley, 2008; Gordon, 2001; Halliwell, 2002; Karadag et al., 2009; Kiokias et al., 2008; Magalhães et al., 2008; Moon and Shibamoto, 2009; McDonald-Wicks et al., 2006; Prior et al., 2005; Roginsky and Lissi, 2005; Laguerre et al., 2007; Sánchez-Moreno, 2002; Singh and Singh, 2008; Strube et al., 1997; Verhagen et al., 2003). The antioxidant activity depends on a number of factors, including type of substrate, medium, and initiators; oxidation conditions, partitioning properties of the antioxidant between lipid and aqueous phases, and the experimental method used (Antolovich et al., 2002; Decker et al., 2005; Frankel and Meyer, 2000; Gordon, 2001; Yanishlieva, 2001). Pitfalls of in vitro antioxidant research include the low stability of antioxidants, which can lead to formation and accumulation of oxidation products (Haenen et al., 2006; Knasmüller et al., 2008). © Woodhead Publishing Limited, 2010
Table 18.1 In vitro methods usually employed for testing the antioxidant activity of pure compounds and extracts. The activities considered include antiradical, anti-lipoperoxidation, metal chelation and reducing capacity Antioxidant activity assay
Substrate/initiator
Scavenging, decolorization assay, indirect competition method of chain lipid peroxidation (chain-breaking)
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bC (b-carotene bleaching)
ROO˙
Scavenging, chain-breaking, competition method
CAT (conjugated autoxidizable triene)
ROO˙
Chain-breaking
Crocin bleaching
ROO˙
Scavenging, chain-breaking, competition method
CUPRAC (cupric ion reducing antioxidant activity)
Cu2+
Reducing capacity 2+
Deoxyguanosine oxidation
HO˙, H2O2, Fe
Deoxyribose oxidation
HO˙, H2O2, Fe2+
Scavenging, chain-breaking, chelating Scavenging, chain-breaking, chelating
DMDP
DMDP˙+ (N,N-dimethyl-pphenylenediamine)
Scavenging, decolorization assay, indirect competition method of chain lipid peroxidation (chain-breaking)
DPPH
DPPH˙ (a,a-diphenyl-bpicrylhydrazyl)
Scavenging, decolorization assay, indirect competition method of chain lipid peroxidation (chain-breaking)
FIC (ferrous ion-chelating assay)
Fe2+, Cu2+
Chelating
FRAP (ferric-reducing antioxidant power)
Fe3+-TPTZ
Reducing capacity, indirect competition method of chain lipid peroxidation (chain-breaking)
FTC (antioxidant activity in the linoleic acid system with ROO˙ ferrothiocyanate reagent)
Scavenging, chain-breaking
HOSC (hydroxyl radical scavenging capacity)
Scavenging, chain-breaking
HO˙
HPS (hydrogen peroxide scavenging)
H2O 2
Scavenging, chain-breaking
Hypochlorous acid scavenging
HOCl
Scavenging, chain-breaking
Extraction of natural antioxidants from plant foods 509
ABTS or TEAC (Trolox equivalent antioxidant capacity) ABTS·+ [2, 2¢-azinobis(3ethylbenzothiozoline-6sulfonate)]
Mechanism
Antioxidant activity assay
Substrate/initiator
Mechanism
Liposome oxidation
ROO˙, Fe2+
Scavenging, chain-breaking, chelating
CD (conjugated dienes)
LDL
Oxidation by copper, aldehydes formation
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In vitro LDL oxidizability
LDL
Oxidation by copper, conjugated dienes determination
Iodometric assay of lipid peroxides
LDL
Iodide oxidation by lipid peroxides
TBARS (thiobarbituric acid-reacting substances)
LDL, ROO˙
Adduct malondialdehyde–thiobarbituric acid (MDA– TBA) formation, scavenging, chain-breaking
ORAC (oxygen-radical absorbance capacity)
Oxygen radical, ROO˙
Scavenging, chain-breaking
Peroxyl radical scavenging
ROO˙
Scavenging, chain-breaking
Peroxynitrite scavenging
ONOO–
Scavenging, chain-breaking
2 –
PHOTOCHEM (photochemiluminescence)
O˙
Scavenging, chain-breaking
PM (phosphomolybdenum method)
Mo6+
Reducing capacity
Reduction of the Fremy’s radical
Fremy’s (potassium nitrosodisulfonate)
Scavenging, indirect competition method of chain lipid peroxidation (chain-breaking)
RP (reducing power)
Fe3+
Reducing capacity
SRS (superoxide radical scavenging)
O 2˙ –
Scavenging, chain-breaking
TAR (total antioxidant reactivity)
ROO˙
Scavenging, indirect competition method of chain lipid peroxidation (chain-breaking)
TOSC (total oxidant scavenging capacity)
ROO˙, O2˙–, HO˙, HOCl, LO(O) , ONOO–, 1O2
Scavenging, chain-breaking
TRAA (a-tocopheroxyl radical attenuating ability)
a-Tocopheroxyl
Scavenging
TRAP (total radical-trapping antioxidant parameter)
ROO˙
Scavenging, chain-breaking
Trichloromethyl peroxyl scavenging
CCl3O2˙
Scavenging
510 Separation, extraction and concentration processes
Table 18.1 Continued
Extraction of natural antioxidants from plant foods 511 The need to use a variety of antioxidant activity assays is generally agreed, as the results obtained in individual tests can be contradictory. Some recommendations on the in vitro determination of antioxidant activity include: identification and determination of active compounds; evaluation of protection against oxidation in foods or physiological model systems under conditions resembling the real chemical, physical, and environmental conditions in the systems to be protected; detailed understanding of the oxidation mechanisms; measurement of both initial and secondary products, and utilization of reagent concentrations that are physiologically relevant. Low oxidation levels should also be determined, and possible mechanism modification under accelerated oxidation conditions should be considered (Antolovich et al., 2002; Choe and Ming, 2006; Collins, 2005; Decker et al., 2005; Frankel, 1993; Frankel and Meyer, 2000; Frankel and Finley, 2008; Gordon, 2001; Prior et al. 2005; Huang et al., 2005; Schlesier et al., 2002). Several classifications of antioxidants assays have been proposed, based on mechanisms such as the ability to quench free radicals by hydrogen donation, the ability for electron transfer, or a combination of both (Huang et al., 2005; Prior et al. 2005; MacDonald-Wicks et al., 2006). Results from simple in vitro experiments are difficult to extrapolate to the heterogeneous conditions of multifaceted in vivo systems (Haenen et al., 2006), where the possible action on the target tissues is affected by absorption, distribution, metabolism and excretion. The animal model used and a battery of well-validated tests to assess the broad diversity of oxidative damage and antioxidative defence parameters, are crucial for antioxidant research in vivo. Measurement of total antioxidant capacity was proposed as an integrated parameter to consider the balanced action of all antioxidants and their synergistic interactions in plasma and body fluids (Aruoma, 2003; Ghiselli et al., 2000; Halliwell, 2009; Sies, 2007). However, pitfalls and serious limits of applicability of this concept have been manifested (Frankel and Finley, 2008; Sies, 2007) as well as the strong dependency on the analytical method used (Fernández-Panchón et al., 2008).
18.3 Natural compounds with antioxidant activity and major sources 18.3.1 Natural compounds Natural antioxidants are synthesized by plants, micro-organisms, fungi, and animals. The most important groups of natural antioxidants are listed below. Phenolics refer to monomeric, oligomeric or polymeric compounds with an aromatic ring bearing one or more hydroxyl substituents and functional derivatives, such as esters, methyl ethers and glycosides. Phenolics (including
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512 Separation, extraction and concentration processes simple phenols, coumarins, flavonoids, stilbenes, lignans, hydrolyzable and condensed tannins, and phlorotannins), have powerful antioxidant activities in vitro (von Gadow et al., 1997; Miller and Ruiz-Larrea, 2002; Tabart et al., 2009) and play a role in signals between plants, plant defense against predation (by micro-organisms, insects and herbivores) or in response to environmental stress (such as air pollution, heavy metal ions and UV-B radiation). Simple phenolics include hydroxybenzoic and hydroxycinnamic acids. The main subclasses of flavonoids are anthocyanins, flavanols, flavanones, flavonols, flavones and isoflavones, and anthocyanidins and their glycosides. Tannins are polyphenolic compounds with varying molecular masses. Plant tannins include hydrolyzable tannins (gallotannins, which yield glucose and gallic acid upon hydrolysis, or ellagitannins, which produce ellagic acid), and condensed tannins (proanthocyanidins). Algae contain phlorotannins, formed by the polymerization of phloroglucinol (Singh and Bharate, 2006). Table 18.2 summarizes the phenolic composition of some plants. Terpenes are a large and diverse class of lipophilic secondary plant metabolites made up of isoprene units, and are classified into hemi-, mono(C10), sesqui- (C15), di- (C20), sester- (C25), tri- (C30), and tetraterpenoids (carotenoids), having eight isoprenoid C 5 residues. Monoterpenes, sesquiterpenes and diterpenes are the main components of essential oils, which also contain oxygenated derivatives and other compounds (including aldehydes, ketones, phenolic, acetates and oxides). Antioxidant potential has been reported for terpenes (Escuder et al., 2002; Grassmann, 2005; Grassmann et al., 2002) as well as synergistic effects with phenolics (Milde et al., 2004). Carotenoids (carotenes and xanthophylls) have received attention because of their provitamin and antioxidant roles (Bohm et al., 2002). Vitamin E includes a family of tocopherols and tocotrienols and some of their ester derivatives. Tissues and seeds of plants are the major sources of vitamin E, which is associated with membrane lipids or appears in lipid storage stuctures. Vitamin E is the most important natural antioxidant in vegetable oil-derived foods from rice bran, palm, and wheat germ (Weber and Rimbach, 2002). The richest source is a by-product of soybean processing (‘oil deodorizer distillate’). Vitamin E effectively inhibits the peroxidation of lipids by peroxyl radical scavenging, and shows relevant activities related to the regulation of enzymes and gene expression (Brigelius-Flohé, 2009). Small proteins and peptides isolated from various protein hydrolyzates show antioxidant activity (Chen et al., 2006; Kitts, 2005; Kitts and Weiler, 2003). Their low-molecular-weight (low-MW) and simple structure, easy absorption, stability under different conditions and the lack of immunoreaction have been cited as advantages (Chen et al., 1996). Some peptides from hydrolyzed food proteins exert antioxidant activities against enzymatic and non-enzymatic peroxidation of lipids and essential fatty acids (Chen et al., 2006; Hook et al., 2001). Certain hydrophobic amino acids and their derivatives present high activity, and show synergistic effects with non-peptidic antioxidants
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Extraction of natural antioxidants from plant foods 513 Table 18.2 Phenolic contents (mg g–1 dry weight) of some plants Plant (Latin name)
Part Phenolic compounds
Achyrocline F satureioides Anethum graveolens F Averrhoa carambola L B Capparis spinosa L
CA, 4CQA, 5CQA, 3,4dCQA, 3,5dCQA, 4,5dCQA pAA, Ap, BA, CA, C, EC, CGA, pCouA, PCys, GA, GeA, K, Lt, Myr, Q, SiA CGA, GA, Q CGA, GA CA, pCouA, FeA, Q, R, VA
TPC/TF
145/48** 78.3 30.1 15.9
Cassine orientalis
L
Castanea vulgaris
L
C, EC, EGC, PCyB2, GA, K, 24/24 Myr, Q Ap, FeA, GeA, Ng, Q, R 11.9
Casuarina equisetifolia Cinnamomun zeylanicum Coffea macrocarpa
B L B
CA, FeA, GA, Q K, Q, R CA, FeA, GA, Q
72.1 57.7 48.9
L
24/18
Diospyros neraudii
L
Eugenia tinifolia
L
Ficus microcarpa
B
Geranium purpureum Ilex paraguariensis
L
Monimiastrum acutisepalum Moringa oliefera
L L
C, EC, ECG, EGC, EGCG, PCyB1, PCyB2, GA, K, Q C, EC, ECG, EGCG, PCyB1, K, Q C, EC, EGC,EGCG, PCyB1, PCyB2, GA, Q, K Cat, Cou, pPGu, pVGu, OlA, pPPh, PCA, PR, Syr, Sy, V, iVA HBA, CA, C, EC, pCouA, GeA, Q, R, SyA, VA CA, 4CQA, 5CQA, 3,4dCQA, 3,5dCQA, 4,5dCQA, 3FQA, 4FQA, 5FQA C, EC, ECG, EGC, PCyB1, GA,Q, K CGA, EA, FeA, GA, Q
Myonima obovata
L
Phytolacca americana Pimpinella anisum
L
L
S
C, ECG, EGC, PCyB1, PCyB2, K, Q HBA, CA, pCouA, FeA, R, VA 4CQA, 5CQA, 3,4dCQA, 3,5dCQA, 4,5dCQA, 3FQA, 4FQA, 5FQA
75/19 23a/17a 237/6.3*
Reference Marques and Farah, 2009 Shyu et al., 2009 Prakash et al., 2007b Proestos et al., 2006 Soobrattee et al., 2008 Proestos et al., 2006 Prakash et al., 2007b Prakash et al., 2007b Soobrattee et al., 2008 Soobrattee et al., 2008 Neergheen et al., 2006 Ao et al., 2008
28.2
Proestos et al., 2006 Marques and Farah, 2009
10a/9a
Neergheen et al., 2006 Prakash et al., 2007b Neergheen et al., 2006 Proestos et al., 2006 Marques and Farah, 2009
32.9 31a/32a 9.2
© Woodhead Publishing Limited, 2010
514 Separation, extraction and concentration processes Table 18.2 Continued Plant (Latin name)
Part Phenolic compounds
Ricinus communis
L
EC, EA, GA, GeA, Q, R
Rubus ulmifolius Ruta graveolens
L
Salix aegyptiaca
C
CA, 4CQA, 5CQA, CGA, FeA, KGu, KCouGPy, KCGPy, QGu HBA, CA, C, GeA, FeA, Q, R CA, C, EGCG, pCouA, GA, R, V CA, C, EGCG, pCouA, GA, Myr, V CA, C, EGCG, pCouA, GA, Myr, Q, R, V CA, pCouA, GeA, Lt, Q, VA
L B Spartium junceum
F
Styrax officinalis
L
Syzygium glomeratum Juglans regia
L
Achillea millefolium
TPC/TF
L H L ST
Reference Singh et al., 2009 Dall’Acqua et al., 2008
4.3
Proestos et al., 2006 107/351** Enayat and Banerjee, 2009 64/165** 212/479** 4.8
HBA, CA, C, EC, CouA, GA, 18.4 GeA, FeA, Ng, Q, VA C, EC, ECG, EGCG, EGC, 84a/39a PCyB1, PCyB2, GA, Q, K 5CQA, 3CouQA, 4CouQA, CouA, QA, QGal, QP-d, QR, QX Ap, Apg, CGA, Lt, Ltdg, Ltg, R, Vi Ap, Apg, CGA, Lt, Ltdg, Ltg, R, Vi Ap, Apg, CGA, Lt, Ltdg, Ltg, R, Vi
Proestos et al., 2006 Proestos et al., 2006 Neergheen et al., 2006 Pereira et al., 2007 Benetis et al., 2008
B, bark; C, catkins; F, flowers; H, herbs; L, leaves; S, seeds; ST, stems. a Composition of phenols (mg g–1 fresh weight) of some plants. TF, total flavonoid expressed as mg quercetin equivalent g–1 on a dry weight basis. * Total flavonoid expressed as mg rutin equivalent g–1 on a dry weight basis. ** Total flavonoid expressed as mg catechin equivalent g–1 on a dry weight basis. TPC: total phenolic content expressed as mg gallic acid equivalent g –1 on a dry weight basis. 3,4dCQA, 3,4-dicaffeoylquinic acid; 3,5dCQA, 3,5-dicaffeoylquinic acid; 3CouQA, 3-coumaroylquinic acid; 3FQA, 3-feruloylquinic acid; 4,5dCQA, 4,5-dicaffeoylquinic acid; 4CouQA, 4-coumaroylquinic acid; 4CQA, 4-caffeoylquinic acid; 4FQA, 4-feruloylquinic acid; 5CQA, 5-caffeoylquinic acid; 5FQA, 5-feruloylquinic acid; Ap, apigenin; Apg, apigenin-7-O-glycoside; BA, benzoic acid; C, catechin; CA, caffeic acid; Cat, catechol; CGA, chlorogenic acid; Cou, coumaran; EA, ellagic acid; EC, epicatechin; ECG, epicatechin-3-gallate; EGC, epigallocatechin; EGCG, epigallocatechin-3-gallate; FeA, ferulic acid; GA, gallic acid; GeA, gentisic acid; HBA, p-hydroxybenzoic acid; iVA, isovanillic acid; K, kaempferol; KCouGPy, kaempferol-3-O-(6≤-p-coumaroyl)-b-d-glucopyranoside; KCGPy, kaempferol-3-O-(6≤-caffeoyl)-b-d-glucopyranoside; KGu, kaempferol-3-O-glucuronide; Lt, luteolin; Ltdg, luteolin-3¢,7-di-O-glycoside; Ltg, luteolin-7-O-glycoside; Myr, myricetin; Ng, naringenin; OlA, oleanolic acid; pAA, p-anisic acid; PCA, protocatechuic acid; pCouA, p-coumaric acid; PCyB1, procyanidin B1; PCyB2, procyanidin B2; PCys, proanthocyanidins; pPGu, p-propylguaiacol; pPPh, p-propylphenol; PR, 4-n-propylresorcinol; pVGu, p-vinylguaiacol; Q, quercetin; QA, quercetin 3-arabinoside; QGal, quercetin 3-galactoside; QGu, quercetin-3-O-glucuronide; QP-d, quercetin-3pentoside derivate; QR, quercetin 3-rhamnoside; QX, quercetin 3-xyloside; R, rutin; SiA, sinapic acid; Sy, syringol; SyA, syringic acid; Syr, syringaldehyde; V, vanillin; VA, vanillic acid; Vi, vicenin-2.
© Woodhead Publishing Limited, 2010
Extraction of natural antioxidants from plant foods 515 such as phenolic compounds (Erdmann et al., 2008). Antioxidant peptides have been obtained from soybean (Pyo and Lee, 2007), soy-fermented foods (Gibbs et al., 2004), soy protein fractions (Moure et al., 2006), barley hordein (Chiue et al., 1997), and potato (Wang and Xiong, 2005). Different hydrolysis conditions result in peptide mixtures with different properties (De Mejia and De Lumen, 2006). Food-derived bioactive peptides from animal or plant proteins have regulatory functions in humans after being released in vitro or in vivo in the gastrointestinal tract (De Mejia and De Lumen, 2006). A number of beneficial health effects have been claimed for food-derived products containing bioactive peptides, including blood pressure-lowering effects, cholesterol-lowering ability, antithrombotic and antioxidant activities, enhancement of mineral absorption and/or bioavailability, and cyto- or immunomodulatory effects. Maillard reaction products (MRP) such as Schiff bases, premelanoidins and melanoidins are formed by reactions involving the condensation of the carbonyl group of reducing sugars with the amino group of amino acids and proteins. Polyphenols, ascorbic acid and other carbonyl compounds can participate in the Maillard reaction. MRP are naturally formed during food processing and storage. Their characteristics and MW depend on both the type of reagents and the processing conditions (temperature, time, pH, water activity) (Jing and Kitts, 2004). The antioxidant properties of MRP are positively correlated with color when the mechanisms responsible for the formation of antioxidants and color follow the same pathway (Manzocco et al., 2001). Formation of potentially harmful substances (acrylamide and hydroxymethylfurfural) increases rapidly with temperature and time (Morales et al., 2009). The antioxidant capacity of MRP formed from model compounds has been reported (Borrelli et al., 2003; Chen and Kitts, 2008; Dittrich et al., 2003; Kim and Lee, 2008; Morales and Babbel, 2002). Maillardderived antioxidants increase the plasma tocopherol, decrease the plasma thiobarbituric acid reactive substances (TBARS) concentration (Somoza, 2005), prevent low density lipoprotein (LDL) oxidation (Mesa et al., 2008), show reducing activity (Yilmaz and Akgun, 2008) and radical scavenging capacity (Michalska et al., 2008), and may be employed as multifunctional ingredients (Lindenmeier et al., 2002). Carbohydrates including polysaccharides, such as dextran, pullulan, mannan, and lipopolysaccharide, exhibit higher free radical scavenging activity than their constituent sugars, although lower than Trolox™ (Tsiapali et al., 2001). Although glucans are weak free radical scavengers in solution, they stimulate free radical activity in a murine macrophage cell line and modestly augment the generation of free radicals. Arabinoglucogalactan from Panax noto ginseng roots exhibited a, a-diphenyl-b-picrylhydrazyl (DPPH) free radical scavenging activity (Wu and Wang, 2008), whereas porphyran, a sulfated polysaccharide isolated from Porphyra, delayed the aging process in mice by enhancing the amount of antioxidative enzymes and reducing the risk of lipid peroxidation (Zhang et al., 2003). Fucoidans,
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516 Separation, extraction and concentration processes sulfated algal polysaccharides made up of fucose as the major component, exhibit antioxidant and biological properties (anticoagulant, anti-inflamatory, antiviral, or antitumoral activities) (Zhao et al., 2008). 18.3.2 Sources The number of potential sources of antioxidant compounds is increasing. Table 18.3 summarizes the composition of some examples of the major vegetal sources of bioactive compounds. Food by-products have the additional advantages of being renewable, widely distributed, largely available and inexpensive. Waste streams from the processing of agricultural or industrial feedstocks are particularly attractive as sources of antioxidants, particularly when they are concentrated and create environmental problems (Balasundram et al., 2006; Moure et al., 2001; Obied et al., 2005). Typical wastes employed for antioxidant production and composition of some representative cases are summarized in Tables 18.4 and 18.5. Table 18.3 Phenolic composition of fruits, vegetables and cereals (% oven-dried substrate) Vegetal source Fruits Acerola (Malpighia emarginata DC.) Andean blackberry (Rubus glaucus Berth.) Blueberry (Vaccinium species) Brazilian mango (Mangifera indica L.) Capulí cherry (Prunus serotina Ehrh. var. Capulí). Grape (Vitis vinifera L.)
Phenolic compounds
Reference
CGA, EC, EGCG, PCyB1, R
Mezadri et al., 2008 pCouA-d, CyG, EA-d, EC, Et, GA, Gs, Vasco et al., PCYs, PeG, Q-d 2009 pHBA, C, CA, CGA, pCouA, Cy, De, Taruscio et al., EC, FeA, Myr, Ma, Peo, Pet, Q 2004 KG,Ma, iMa, Q, QAp, QAf, QG, QGal, Ribeiro et al., QR, QX 2008 C, CGA, CyG, EC, K-d, PCYs, Q-d Vasco et al., 2009
C, CA, HCA, oCouA, pCouA, GA, R, Q Kedage et al., 2007 Lemon (Melissa officinalis CA, mCouA, Er7G, He, Het, Na, Ng, Dastmalchi L.) RA et al., 2008 Plum (Prunus salicina CA-d, nCGA, pCouA-d, Cyg, EC, K-d, Vasco et al., Lindl.) (var. ‘Santa Rosa’ PCYs, Q-d 2009 and ‘Beauty’) Portuguese pear (Pyrus Ar, monC, terC, CQA, pCouMA, Ferreira et al., communis L. var. S. monEC, terEC, extEC 2002 Bartolomeu) Strawberry (Fragaria × CyG, EA, K, K-CouG, Pe, PeG, PeR, Zhang et al., ananassa Duch.) HPhAC, Q 2008 CGA, nCGA, pCouQA, CyG, EC, PeR, Usenik et al., Sweet cherry (Prunus avium L.) PeoR, R 2008 © Woodhead Publishing Limited, 2010
Extraction of natural antioxidants from plant foods 517 Table 18.3 Continued Vegetal source
Phenolic compounds
Reference
Vegetables Broccoli (Brassica oleracea L.) Escarole (Cichorium indivia var. latifolium L.) Lettuce (Lactuca sativa var. capitata L.) Onion (Allium cepa)
nCGA, FeGe, PCA, dSiGe, SiFeGe, SiGe, dSiFeGe, SidFeGe dCQA, CGA, CGA-d, CtA, ChiA, KGl, KMG dCQA,CGA, CGA-d, CtA, ChiA, Q-d, QMG FeA, GA, K, PCA, Q
Vallejo et al., 2003 Degl’innoocenti et al., 2008 Degl’innoocenti et al., 2008 Prakash et al., 2007 Reddivari et al., 2007 Degl’innoocenti et al., 2008 Aehle et al., 2004
Potato (Solanum Axa, C, CA, CGA, Cxa, GA, Lu, Vxa tuberosum L.) Rocket salad (Eruca sativa HCAs-d, K-d Mill.) Spinach (Spinacia Pa, Sp oleracea) Cereals Barley (Hordeum vulgare C, CA, pCouA, EC, FeA, GA, PCA, L.) SyA, VA Corn (Zea mays L.) CyG, aCyG, HCAs, pCouA, PCA, He-d, PeG, aPeG, PeoG, aPeoG, Q-d, VA Millets
pHBA, C, CA, CinA, pCouA, FeA, GA,GeA, L, Or, iOr, PCA, SaA, Sap, SiA, Tr, VA, Vi, Vt, iVt Rice (Oryza sativa L.) HBA, CA, CGA, pCouA, FeA, FS, PCA, SiA, SiS, SyA, VA Rye (Secale cereale L.) ALRs, pHBA, CA, FeA, SiA, SyA, VA, VeA Sorghum (Sorghum Apf, Ap, Api, ApiG, MApi, MApiG, bicolor L.) 5MApi, PApi, pHBA, CA, EC, pCouA, CinA, PCyB1, PDe, Er, Er5G, FeA, GA, PCA, GeA, KRG, Lut, Ltn, LtnG, PLtn, Lt, MLtn, MLtnG, 7MLtn, Ng, Ta, TaG, SaA, SiA, SyA, VA Wheat (Triticum aestivum CA, oCouA, pCouA, FeA, SyA, VA L.)
Zhao et al., 2006 Pedreschi and CisnerosZevallos, 2007 Dykes and Rooney, 2006 Tian et al., 2004 Heiniö et al., 2008 Dykes and Rooney, 2006
Mpofu et al., 2006
5MApi, 5-methoxyapigeninidin; 7-MApiG, 7-methoxyapigeninidin-5-glucoside; 7MLtn, 7-methoxyluteolinidin; aCyG, acylated cyanidin-3-glucoside; ALRs, alkylresorcinols; Ap, apigenin; aPeG, acylated pelargonidin-3-glucoside; aPeoG, acylated peonidin 3-glucoside; Apf, apiforol; Api, apigeninidin; ApiG, apigeninidin-5-glu; Ar, arbutin; Axa, antheraxanthin; C, catechin; CA, caffeic acid; CA-d, caffeic acid derivatives; CGA, chlorogenic acid; CGA-d, chlorogenic acid derivates; ChiA, chicoric acid; CinA, cinnamic acid; CQA, caffeoylquinic acid; CtA, caffeoyltartaric acid; Cxa, canthaxanthin; Cy, cyanidin; Cyg, cyanidin glycosides; CyG, cyanidin-3-glucoside; dCQA, dicaffeoylquinic acid; De, delphinidin; dSiFeGe, 1,2¢-disinapoyl-2-feruloylgentiobiose; dSiGe, 1,2-disinapoylgentiobiose; EA, ellagic acid; EA-d, ellagic acid derivatives; EC, epicatechin; EGCG, epigallocatechin gallate; Er, eriodictyol; Er5G, eriodictyol 5-glucoside; Er7G, eriodictyol-
© Woodhead Publishing Limited, 2010
518 Separation, extraction and concentration processes Table 18.3 Continued 7-O-glucoside; Et, ellagitannins; extEC, extension epicatechin; FeA, ferulic acid; FeGe, 1,2diferuloylgentiobiose; FS, 6¢-O-feruloylsucrose; GA, gallic acid; GeA, gentisic acid; Gs, galloyls; HBA, hydroxybenzoic acid; HCA, hydrocaffeic acid; HCAs, hydroxycinnamic acids; HCAs-d, hydroxycinnamic derivates; He, hesperidin; He-d, hesperidin derivates; Het, hesperetin; HPhAC, 3,4,5trihydroxyphenyl-acrylic acid; iMa, isomangiferin; iOr, isoorientin; iVt, isovitexin; K, kaempferol; K-CouG, kaempferol-3-(6¢-coumaroyl)glucoside; K-d, kaempferol derivatives; KG, kaempferol 3-O-glucoside; KGl, kaempferol-3-O-glucuronide; KMG, kaempferol-3-O-(6-O-malonylglucoside); KRG, kaempferol 3-rutinoside-7-glucuronide; L, lucenin-1; Lt, luteolin; Ltn, luteolinidin; LtnG, luteolinidin-5-glucoside; Lu, lutein; Lut, luteoforol; Ma, mangiferin; MApi, 7-methoxyapigeninidin; mCouA, m-coumaric acid; MLtn, 5-methoxyluteolinidin; MLtnG, 5-methoxyluteolinidin-7-glucoside; monC, monomeric catechin; monEC, monomeric epicatechin; Myr, myricetin; Na, naringin; nCGA, neochlorogenic acid; Ng, naringenin; oCouA, o-coumaric acid; Or, orientin; Pa, patuletin; PApi, proapigeninidin; PCA, protocatechuic acid; pCouA, p-coumaric acid; pCouA-d, p-coumaric acid derivatives; pCouMA, p-coumarylmalic acid; pCouQA, p-coumaroylquinic acid; PCyB1, procyanidin B1; PCYs proanthocyanidins; PDe, prodelphinidin; Pe, pelargonidin; PeG, pelargonidin-3-glucoside; Peo, peonidin; PeoG, peonidin 3-glucoside; PeoR, peonidin 3-rutinoside; PeR, pelargonidin-3rutinoside; Pet, petunidin; pHBA, p-hydroxybenzoic acid; PLtn, proluteolinidin; Q, quercetin; QAf, quercetin 3-O-arabinofuranoside; QAp, quercetin 3-O-arabinopyranoside; Q-d, quercetin derivatives; QG, quercetin 3-O-glucoside; QGal, quercetin 3-O-galactoside; QMG, quercetin-3-O(6-O-malonylglucoside); QR, quercetin 3-O-rhamnoside; QX, quercetin 3-O-xyloside; R, rutin; RA, rosmarinic acid; SaA, salicylic acid; Sap, saponarin; SiA, sinapic acid; SidFeGe, 1-sinapoyl-2,2¢diferuloylgentiobiose; SiFeGe, 1-sinapoyl-2-feruloylgentiobiose; SiGe, 1,2,2¢-trisinapoylgentiobiose; SiS, 6¢-O-sinapoylsucrose; Sp, spinacetin; SyA, syringic acid; Ta, taxifolin; TaG, taxifolin-7-glucoside; terC, terminal catechin; terEC, terminal epicatechin; Tr, tricin; VA, vanillic acid; VeA, veratric acid; Vi, violanthin; Vt, vitexin; Vxa, violaxanthin.
Table 18.4 Residual sources proposed for antioxidant manufacture Feedstock
References
Residues from grape and wine
Cruz et al., 2004; Corrales et al., 2008; Lafka et al., 2007; Makris et al., 2007; Pinelo et al., 2006; Arvanitoyannis et al., 2006; Louli et al., 2004
Apple pomace
Schieber et al., 2003
Olive
Obied et al., 2005; McDonald et al., 2001
Husks and hulls
Moure et al., 2000; Conde et al., 2008; Goli et al., 2005; Takeoka and Dao, 2003; Oliveira et al., 2008; DeliormanOrhan et al., 2009; Rubilar et al., 2007
Woods
Pérez-Bonilla et al., 2006; Castro et al., 2008; Moure et al., 2005
Leaves
Ferreira et al., 2007; Ozsoy et al., 2008
Seeds
Liu and Yao, 2007; Siddhuraju and Becker, 2007
Peels
Berardini et al., 2005; Li et al., 2006
Fruits and vegetables and their processed products and by-products have been extensively investigated. Their activity has been ascribed to the presence of phenolic compounds (Cieślik et al., 2006; Heinonen and Meyer, 2002; Macheix et al., 1990; Peschel et al., 2006; Proteggente et al., 2003; Robards et al., 1999; Sakakibara et al., 2003). Abundant research on fruits © Woodhead Publishing Limited, 2010
Extraction of natural antioxidants from plant foods 519 Table 18.5 Chemical composition of solid agro-industrial by-products (% oven-dried substrate) Solid agro-industrial Phenolic compounds by-product Almond hulls
References
CGA, cCGA, nCGA
Takeoka and Dao, 2003 Almond shells HCinAs, pCouA, VA, SyA Moure et al., 2007 Apple, peach and CA, p-CouA, FeA Gorinstein et al., pear peels 2002 Apple peels CA, C, EC, CGA, FeA, Q, QA, QG, QGal He and Liu, 2008 Schieber et al., 2003 Apple pomace C, EC, CGA, pCouA, pCouQA, PCyB2, FeA, Pht, Phx, Phz, Q, QA, QG, QGal, QR, QRu, QX Banana peel/pulp C, EC, CG Someya et al., 2002 Barley husks dHB, pCouA, GA, SyA, VA, V Conde et al., 2008 Quettier-Deleu Buckwheat hulls/ B2D, B2G, EC, ECG, QGal, QRu, Q et al., 2000 flour Chestnut wood 4HB, Con, EA, FeA, GA, Sin, Sc, SyA, Canas et al., 1999 Syra, V, VA, Um Coconut husks 4HBA, FeA Dey et al., 2003 Corn cobs AV, HMCin, iEu, MEu, Gu, Gua, EGu, Garrote et al., 2007a VGu, VPh, Sy, ASy, Syra, V Citrus peel and seed CA, p-CouA, Er, nEr, FeA, He, nHe, Na, Bocco et al., 1998 Nr, SiA Eucalypt wood 3HB, HMCin, Co, dHEu, MEu, Gu, VGu, Garrote et al., 2007b iEu, V, iVA, HVA Grape skin and seeds C, EC, CA, oCouA, pCouA, CyGl, GA, Lafka et al., 2007 PCA, SyA, Ty, HTy, VA Olive tree leaves ApG, C, CA, Di, DiG, Lt, LtG, Ol, R, Ty, Benavente-García HTy, V, VA, Ve et al., 2000 Olive tree pruning dHB, HMCin, Ol, Sy, Syra, Ty, HTy, V, Conde et al., 2009 HV, VA Potato peel GA, CGA, CA Kanatt et al., 2005 Rice husks AV, 4HB, HMCin, VPh, Gu, Gua, VGu, Garrote et al., 2007 Eu, iEu, MEu, Sy, ASy, Syra, V, HVA Sweetpotato leaves CA, CGA, 3,4dCQA, 3,5dCQA, Islam et al., 2002 4,5dCQA, tCQA Winemaking waste dHB, 3HBA, CinA, pCouA, GA, PCA, Cruz et al., 2004 solids SyA, VA, iVA Wood of olive tree De, iEu, MEu, Gu, Sy, Syra, HSyA, Ty, Castro et al., 2008 branches HTy, Ol, V, HV 3,4dCQA, 3,4-di-O-caffeoylquinic acid; 3,5dCQA, 3,5-di-O-caffeoylquinic acid; 3HB, 3-hydroxybenzaldehyde; 3HBA, 3-hydroxybenzoic acid; 4,5dCQA, 4,5-di-O-caffeoylquinic acid; 4-HB, 4-hydroxybenzaldehyde; 4HBA, 4-hydroxybenzoic acid; ApG, apigenin-7-glucoside; ASy, acetosyringone; AV, acetovanillone; B2D, proanthocyanidin B2; B2G, B2-3´-O-gallate; C, catechin; CA, caffeic acid; cCGA, cryptochlorogenic acid; CG, catechin gallate; CGA, chlorogenic acid; CinA,
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520 Separation, extraction and concentration processes Table 18.5 Continued cinnamic acid; Co, coniferol; Con, coniferaldehyde; CYGl, cyanidin glycosides; De, desaspidinol; dHB, 3,4-dihydroxybenzaldehyde; dHEu, dihydroeugenol; Di, diosmetin; DiG, diosmetin-7-glucoside; EA, ellagic acid, EC, epicatechin; ECG, epicatechin gallate; EGu, 4-ethylguaiacol; ER, eriocitrin; Eu, 4-eugenol; FeA, ferulic acid; GA, gallic acid; Gu, guaiacol; Gua, guaiacylacetone; HCinAs, hydroxycinnamic acids; He, hesperidin; HMCin, 4-hydroxy-2-methoxycinnamaldehyde; HSyA, homosyringic acid; HTy, hydroxytyrosol; HV, homovanillyl alcohol; HVA, homovanillic acid; iEu, isoeugenol; iVA, isovanillic acid; Lt, luteolin; LtG, luteolin-7-glucoside; MEu, methoxyeugenol; Na, naringin; nCGA, neochlorogenic acid; nEr, neoeriocitrin; nHe, neohesperidin; Nr, narirutin; oCouA, o-coumaric acid; OL, oleuropein; PCA, protocatechuic acid; pCouA, p-coumaric acid; pCouQA, p-coumaroylquinic acid; PCY-B2, procyanidin B2; Pht, phloretin; Phx, phloretin xyloglucoside; Phz, phloridzin; Q, quercetin; QA, quercetin-3-arabinoside; QG, quercetin-3-glucoside; QGal, quercetin3-galactoside; QR, quercetin-3-rhamnoside; QRu, quercetin-3-rutinoside; QX, quercetin-3-xyloside; R, rutin; Sc, scopoletin; SiA, sinapic acid; Sin, sinapaldehyde; Sy, syringol; SyA, syringic acid; Syra, syringaldehyde; tCQA, 3,4,5-tri-O-caffeoylquinic acid; Ty, tyrosol; Um, umbelliferone; V, vanillin; VA, vanillic acid; Ve, verbascoside;VGu, 4-vinylguaiacol; VPh, 4-vinylphenol.
is available, including citrus (Dastmalchi et al., 2008; Di Majo et al., 2005; Jayaprakasha and Patil, 2007; Tripoli et al., 2007), berry fruits (Szajdek and Borowska, 2008; Vasco et al., 2009; Taruscio et al., 2004; Zhang et al., 2008) and cherries (Piccolella et al., 2008). Results have also been reported for juices (Gardner et al., 2000), teas (Gupta et al., 2008; von Gadow et al., 1997; Wiseman et al., 1997) coffee (Nardini et al., 2002), wines (Burns et al., 2001; Fogliano et al., 1999; Minussi et al., 2003; Sanchez-Moreno et al., 2003; Villaño et al., 2005), and several beverages (Arts et al., 2000; Richelle et al., 2001; Tabart et al., 2009). The potential of aromatic herbs, spices and medicinal plants was first assessed in the 1950s (Chipault et al., 1952). Their major antioxidant components are vitamins, phenolic acids, flavonoids and flavonoid derivatives, terpenoids, carotenoids, phytoestrogens and minerals. Specific compounds from plants include carnosic acid, carnosol, rosmarinic acid, rosmanol, thymol, carvacrol, gingerol-related compounds, curcumins, capsaicin, capsaicinol, and ascorbic acid (Suhaj, 2006; Yanishlieva et al., 2006). Their effectiveness in different substrates and other biological activities have been reviewed (Srinivasan, 2005; Yanishlieva et al., 2006; Zheng and Wang, 2001). The in vitro antioxidant capacity of cereals is significantly correlated with their polyphenol content, except for maize (Decker et al., 2002). During digestion, the antioxidant capacity of cereals is increased, providing a favorable antioxidative environment for the epithelium in the large intestine (Fardet et al., 2008). However, the in vitro antioxidant capacity of cereals is only an approximate reflection of their in vivo antioxidant effect, owing to differences in solubility/bioavailability within the digestive tract and to the metabolism/conjugation with compounds such as polyphenols. About 6000 seaweed species have been identified and they are grouped into green (chlorophytes), brown (pheophytes) and red (rhodophytes) algae.
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Extraction of natural antioxidants from plant foods 521 Seaweeds are traditionally exploited mainly for soluble polysaccharides and direct utilization as foods. The soluble fraction of red seaweeds is mostly composed of sulfated galactans (agar, carrageenans), whereas the brown seaweeds (alginates, fucans, and laminarans) correspond mostly to reserve b-glucans. The recovery of valuable biomolecules, including vitamins, minerals, lipids, polyphenols, proteins, sterols, pigments, tocopherol derivatives and related isoprenoids, is a topic of growing interest (Herrero et al., 2006; Kumar et al., 2008). The most potent antioxidants of algae are phlorotannins and sulfated polysaccharides (Rupérez et al., 2002; Wang et al., 2008; Xue et al., 2001; Ye et al., 2008). A recent work has reported the irradiation-dependent production of butylhydroxytoluene (BHT) by green algae and cyanobacteria (Babu and Wu, 2008). Mushrooms are also a source of phenolic antioxidants (Cheung et al., 2003).
18.4 Biological activities of natural antioxidants Reactive oxygen species (ROS) play a dual role: a certain physiological level of ROS is crucial for the regulation of cell functions (Wang and Yi, 2008), but ROS and other radicals are also involved in diabetes, cancer, liver disease, aging, arthritis, AIDS, macular degeneration, and autoimmune, inflammatory, cardiovascular and neurodegenerative diseases. Antioxidants prevent ROS concentrations from reaching a damaging level and have been considered as a promising therapy for the prevention and treatment of these diseases (Seifried et al., 2007). The results from many intervention trials with antioxidants failed to demonstrate benefits in humans, probably because the reactive species may not be important, the pathologies were too advanced, the doses used may be wrong or the administered antioxidants do not decrease the oxidative damage (Halliwell, 2009). Antioxidant properties alone are not sufficient to explain their biological properties, some of which are indicated in Table 18.6. Although an antioxidant is a redox agent that in the presence of metal ions could act as a pro-oxidant, in vivo most transition metal ions are protein-conjugated and not available to catalyze free radical reactions (Hadi et al., 2007). The most studied compounds for controlling various disorders (including cardiovascular, neurological and neoplastic diseases) are plant-derived polyphenols. The use of the cardioprotective properties of dietary antioxidants against cardiovascular diseases has been studied and some conflicting findings have been found (Seifried et al., 2007). Natural polyphenolic compounds possess antioxidant, vasorelaxant and antihypertensive properties (Depeint et al., 2002; Erlund, 2004; Ghosh, 2005; Ullah and Khan, 2008). Some effects of polyphenols (such as anti-atherogenesis) can be attributed to their antioxidant properties (German and Walzem, 2000). Selected polyphenols
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Biological activities
Plant or extract
Active compounds
Reference
Antiallergic Antiarthritic
Coronopus didymus Cleome gynandra
Antiasthmatic Antibacterial Anticancer Anticoagulant Antidiabetic Antiglycative Antihyperglycemic
Laurencia undulata Cyperus rotundus Prunus serrulata Cirsium japonicum Psidium guajava L. Helichrysum plicatum Psidium guajava L. Punica granatum
Mantena et al., 2005 Narendhirakannan et al., 2005 Jung et al., 2009 Kilani et al., 2008 Lee et al., 2007 Yin et al., 2008 Hsieh et al., 2007 Aslan et al., 2007 Hsieh et al., 2007 Bagri et al., 2009
Antihyperlipidemic
Coronopus didymus Helichrysum plicatum Punica granatum
Flavonoids, tannins Polyphenolic compounds, tannins, anthroquinones, flavonoids Phlorotannins Tannins, flavonoids, coumarins, sterols Phenolic glucosides, flavonoids, anthocyanins Phenolic acids, flavonoids Phenolic compounds Phenolic acids, flavonoids Phenolic compounds Alkaloids, flavonoids, anthocyanins, hydrolyzable tannins Flavonoids, tannins Phenolic acids, flavonoids Alkaloids, flavonoids, anthocyanins, hydrolyzable tannins Flavonoids and phenolics acids Polyphenols, tannins Bioflavonoids, catechins, procyanidins, phenolic acids Monoterpene glycosides Eugenol Curcuminoids Polyphenols Oligomeric procyanidins, catechins, condensed tannins
Souza et al., 2007 Kim et al., 2007 Hasegawa et al., 2008 Panich et al., 2009 Panich et al., 2009 Ahn et al., 2008 Esquenazi et al., 2002
Anti-inflammatory
Antimelanogenic
Antimicrobial
Geranium robertianum L. Uncaria tomentosa Myracrodruon urundeuva Pycnogenol® Eucalyptus globulus Almina galanga Curcuma aromatic Sasa quelpaertensis Nakai Cocos nucifera L.
Mantena et al., 2005 Aslan et al., 2007 Bagri et al., 2009 Amaral et al., 2009
522 Separation, extraction and concentration processes
Table 18.6 Biological activities reported for antioxidant extracts from some plant foods
Pycnogenol® and Ginkgo biloba Hibiscus tiliaceus L.
Antiparkinson Antipyretic Antitumor Antiulcer Antiviral Apoptotic activities Cytotoxic Hepatoprotective
Mucuna pruriens Coronopus didymus Senna alata Myracrodruon urundeuva Cocos nucifera L. Cyperus rotundus Cyperus rotundus Coronopus didymus Commiphora berryi
Polyphenols, procyanidins Polyphenols, carotenoids, tocopherols, flavonoids, anthocyanins, phytosterols Polyphenols Flavonoids, tannins Flavonoids, tannins, polyphenols Polyphenols, tannins Oligomeric procyanidins, catechins, condensed tannins Tannins, flavonoids, coumarins, sterols Tannins, flavonoids, coumarins, sterols Flavonoids, tannins Phenolic compounds, tannins
Decalepis hamiltonii
Phenolic compounds
Ecklonia cava
Polyphenols
Dhanasekaran et al., 2008 Mantena et al., 2005 Pieme et al., 2008 Souza et al., 2007 Esquenazi et al., 2002 Kilani et al., 2008 Kilani et al., 2008 Mantena et al., 2005 Gowri Shankar et al., 2008 Srivastava and Shivanandappa, 2009 Ahn et al., 2008
Morus alba L.
Flavonol glycosides
Katsube et al., 2006
Castanea sativa
Polyphenols
Almeida et al., 2008
Immunomodulatory activity Inhibitor LDL oxidation Skin protector
Križková et al., 2008 Rosa et al., 2006
Extraction of natural antioxidants from plant foods 523
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Antimutagenic
524 Separation, extraction and concentration processes (extracts from grape and wine, resveratrol, genistein, dihydrodaidzeins) relax endothelial function (Shen et al., 2006). Flavonoids can have favorable effects on atherosclerosis. Lipid-lowering activity has been reported for tea flavonoids (Li et al., 2005), a-tocopherol, and b-carotene, which inhibit the oxidation of LDL and the atherosclerotic process (Odeh and Cornish, 1995). Tocopherols show enhanced immune response and regulation of platelet aggregation (Weber and Rimbach, 2002). Protection against coronary heart disease was reported for resveratrol (Szewczuk and Penning, 2004), anthocyanidins, EGCG, and vitamin E (Violi et al., 2006). Biologically active peptides promote diverse activities (antithrombotic, hypocholesterolemic and antihypertensive actions) relevant to cardiovascular health (Erdmann et al., 2008). Inhibitors of angiotensin-converting enzyme (ACE) play an important role in the regulation of blood pressure and fluid and salt balance in mammals. Immunomodulatory activity has also been reported for peptides (Erdmann et al., 2008). Antiallergenic activity has been reported for flavonoids (Das and Rosazza, 2006; Ghosh, 2005; Hodek et al., 2002; McKay and Blumberg, 2006), which may also preserve T-cell-mediated immunity (Horváthová et al., 2001; Strickland, 2001). Bioactive compounds such as flavonoids (Horváthová et al., 2001) and vitamin E protects against cataracts. Vitamin E decreases oxidative stress and the levels of erythrocyte osmotic fragility in patients on dialysis (Uzum et al., 2006) and had beneficial effects on diabetic patients (Levy and Blum, 2007). Epidemiological studies suggested a positive association between a diet rich in fruit and vegetables and a lower incidence of cancer (Ames et al., 1993; Pan et al., 2008a; Ren et al., 2003; Tanaka and Sugie, 2008). The antioxidant properties may only partly explain the antitumor effects of dietary phenolics, which may exert modulatory actions in cells, interfering in the steps leading to the development of malignant tumors (Kanakis et al., 2007). Supplemental antioxidants may help to prevent cancer only in dietdeficient populations or individuals, and antioxidant anticancer agents may affect patients differently according to their health (Vickers, 2007). Although dietary antioxidants have attracted great interest and are considered safe, they have to pass controlled clinical trials for therapeutic or chemopreventive use, and potential toxic flavonoid–drug interactions should be considered (Galati and O’Brien, 2004). Many natural agents have shown potential in bioassays and animal models, and some of them have been selected for ongoing phase I–III clinical trials (Bonfili et al., 2008). An ideal, effective and acceptable agent should be non-toxic in normal and healthy cells, low cost and efficient against multiple cancers, act according to a known mechanism, deserve public acceptance, and be suitable for oral consumption. Potential cancer preventive agents suppress carcinogenesis by several major mechanisms: inhibiting phase I enzymes or blocking carcinogen formation, inducing phase II (detoxification) enzymes, scavenging DNA reactive agents, suppressing the over-expression of pro-oxidant enzymes, modulating hormone homeostasis,
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Extraction of natural antioxidants from plant foods 525 suppressing hyper-cell proliferation induced by carcinogens, inducing apoptosis, counteracting angiogenesis, and/or inhibiting certain phenotypic expressions of preneoplastic and neoplastic cells (D’Archivio et al., 2008; Shklar, 1998; Ren et al., 2003; Pan et al., 2008a). Various dietary antioxidants are under investigation for their anticancer properties, including curcumin (Tsvetkov et al., 2005; Maheshwari et al., 2006), oleuropein (Bonfili et al., 2008), and tangeretin, nobiletin and resveratrol (Narayanan, 2006). Green tea inhibits tumor incidence and multiplicity in various organs, and recent phase I and II clinical trials have been conducted to explore its anticancer effects (Dou, 2008). Flavonoids in human diet may reduce the risk of various cancers (Kanadaswami et al., 2005), including hormone-dependent breast and prostate cancers (Hodek et al., 2002), intestinal neoplasia (Hoensch and Kirch, 2005) and skin cancer (Singh and Agarwal, 2002). Lycopene shows both in vitro and in vivo anticancer activities, possibly through ROS scavenging, up-regulation of detoxification systems, interference with cell proliferation, induction of gap-junctional communication, inhibition of cell cycle progression and modulation of signal transduction pathways (Bhuvaneswari and Nagini, 2005). A beneficial role of vitamin C on cancer has been claimed, although further clinical trials are needed (Verrax and Calderón, 2008). Vitamin E and vitamin E-based compounds are potent antioxidants regulating peroxidation, some forms display apoptotic activity against cancer cells and restore multidrug resistant tumor cells sensitivity to chemotherapeutic agents (Kline et al., 2007; Sylvester, 2007). Among soy peptides, lunasin suppresses transformation of cells induced by carcinogens and viral oncogenes (De Mejia and De Lumen, 2006), and the Kunitz trypsin inhibitor suppresses ovarian cancer cell invasion by blocking urokinase up-regulation (Xiao et al., 2005). Combinations of antioxidant nutrients act synergistically, and their use is proposed to offer a better quality of life (Shklar, 1998), and to present a co-operative action with chemotherapeutic drugs and radiation therapy (Bonfili et al., 2008; Prasad et al., 2002). The nervous system, rich in fatty acids and iron, is vulnerable to free radical generation (Ullah and Khan, 2008). Reactive species are constantly produced in the brain, for example by leakage of electrons from the mitochondrial electron transport chain to generate superoxide radical (O2–). Increased levels of oxidative damage may occur early in the progression of many neurodegenerative diseases (Philips et al., 1993), as oxidative damage may aggravate the accumulation and precipitation of proteins. Epidemiological studies indicate that dietary antioxidants can limit the incidence of neurodegenerative disorders (Morris, 2002). Polyphenols are active agents in neuroprotection owing to their ability to influence and modulate several cellular processes such as signaling, proliferation, apoptosis, redox balance, and differentiation. Their role in neurological disorders has been studied (Dajas et al., 2002; Horváthová et al., 2001; Singh and Singh, 2008). Promising compounds for neuroprotection include polyphenols such as epigallocatechin gallate (EGCG), curcumin, naringenin, extracts of
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526 Separation, extraction and concentration processes blueberries and Scutellaria (Ullah and Khan, 2008), and extracts from green tea (Bastianetto et al., 2006). Flavonoids and their derivatives are reported to inhibit the growth and development of HIV by suppressing acute infection and inhibiting protease, integrase and reverse transcriptase (Critchfield et al., 1996), whereas scutellarin showed anti-HIV activity (Zhang et al., 2005). The overall role of reactive species in chronic inflammatory diseases is not clear. Reactive species may help explain why such diseases increase cancer risk, but ironically these species can sometimes be anti-inflammatory. Hence the use of antioxidants to treat chronic inflammatory diseases may not be as simple as it originally sounded (Halliwell, 2009). Natural antioxidants with anti-inflamatory activity include phenolics (Lomnitski et al., 2000), resveratrol, flavonoids (Ghosh, 2005; Das and Rosazza, 2006; Hodek et al., 2002; Kim et al., 2004), terpenoids (De las Heras et al., 2003), and glucans (Tsiapali et al., 2001). Therapeutic potential for combating bronchial asthma has been reported for polyphenolic extracts from Laurencia undulata, an edible red alga (Jung et al., 2009). Tyrosinase inhibitors can ameliorate cutaneous hyperpigmentary disorders. Their use is becoming increasingly important in the cosmetic industry owing to their skin whitening action and preventive effects, and there is a growing interest in their medicinal and cosmetic applications (Parvez et al., 2007). Some carotenoids (a- and b-carotene and b-cryptoxanthin) are precursors of vitamin A and protect against chemical oxidative damage and against several kinds of cancer, age-related macular degeneration and UV-induced erythema.
18.5 Extraction of natural antioxidants from plant foods and residues Solid–liquid extraction is a heterogeneous operation involving transfer of solutes from a solid to a fluid. Extractable compounds of vegetal origin are a complex mixture of solutes that can be extracted at different rates depending on their location (outer surface, pores and vacuoles). Extraction involves the following sequential steps: (i) transport of solvent from the bulk solution to the external surface of the particle, (ii) solvent penetration and diffusion in the solid matrix, (iii) solubilization of the components, (iv) transport of the solute(s) through the solid matrix, and (v) transport of the solute(s) from the external surface of the solid to the bulk solution. 18.5.1 Conventional solvent extraction The type of solvent is one of the most influential variables on both extraction yield and type of extracted solutes. Methanol, ethanol, and water are widely
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Extraction of natural antioxidants from plant foods 527 employed for extracting phenols. Methanol exhibits the highest capacity to extract phenolics (Pinelo et al., 2005), but shows toxicity, whereas hot water may extract polyphenols without safety concerns. Ethanol–water mixtures are suited to penetrate the hydrophobic areas of the vegetable matrix and help to precipitate soluble proteins, facilitating further processing. Representative examples of extraction processes involving these solvents are summarized in Table 18.7. Temperature is one of the most influential variables affecting the release of phenols. Increased temperatures favor extraction by enhancing both solubility and diffusion. However, temperature cannot be increased indefinitely, because instability of phenolic compounds and denaturation of membranes may take place at temperatures above 50 °C (Cacace and Mazza, 2003). Partially oxidized polyphenols can exhibit higher antioxidant activity than non-oxidized phenols. Thermal treatment of crude extracts can be performed to modify their composition by oxidation, hydrolysis and isomerization, and can result in the formation of compounds with new antioxidant properties. Therefore, heat-induced antioxidants balance the thermal loss of antioxidants, or the overall antioxidant properties can be increased upon heating (Delgado-Andrade and Morales, 2005). This behavior was reported for catechin, resveratrol, grape extract (Pinelo et al., 2005), and citrus peel (Xu et al., 2007). Steam-processing of broccoli may release more bound phenolic acids from the breakdown of cellular constituents (Roy et al., 2009). Heating vegetal products such as sweet corn increased the total antioxidant activity and the phenolic content (Dewanto et al., 2002). Heat treatment of huyou (Citrus paradisi Changshanhuyou) peel increased the free fraction of phenolic acids and decreased the ester, glycoside, and ester-bound fractions (Xu et al., 2007), whereas the antioxidant capacity of citrus peel [measured by the ABTS and ferric-reducing antioxidant power (FRAP) assays] increased with heating time and temperature owing to the increase of lower MW phenolics (benzoic acids and cinnamic acids). High solvent-to-solid ratios favor phenol yields, but a balance between high costs (derived from solvent usage) and absence of saturation effects has to be found (Pinelo et al., 2005). However, operational variables such as temperature, solvent-to-solid ratio, and type of solvent used only affect solutes weakly linked to the cell wall structure and those contained in vacuoles. Extraction of compounds forming part of the cell walls may require other strategies, such as cell wall breaking or enzymatic treatments. Even though conventional extraction with organic solvents is widely used, it has low efficiency and requires large volumes of solvent, long extraction times and high temperature. This could lead to loss of biological activities and decreased selectivity, as well as to problems derived from organic solvent cost and disposal (Robards 2003). Additionally, organic solvents must be completely removed from the exhausted solids. Recently, there has been an increasing interest in using environmentally friendly extraction technologies for producing extracts of high quality and activity. Some of these techniques
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Raw material
Solvent
Solubles yield or Antioxidant activity TEP* (% dry weight)
Reference
Anethum graveolens (flowers)
W
8
DPPH(–)*1; bC(–)*1
Sharififar et al., 2009
Anethum graveolens (flowers)
E
2.84
TEAC(–)*2; DPPH(–)*2,3; RP(–)*2,3; FIC(–)*4; bC(–)*2
EA
8.28
Hordeum vulgare (seeds) Chamaecyparis obtusa (bark)
E (70%)
4.70
TEAC(–) ; DPPH(–) ; RP(–) ; FIC(–) ; bC(–) RP(–)*1,3; DPPH(–)*1,3; FTC(+)*1(=)*3
E
9.07 (GA)
DPPH(–)*2; TEAC(–)*5; RP(–)*5; bC(–)*5
EA
10.28 (GA)
Cirsium japonicum (leaves) Chukrasia tabularis (leaves)
W
24.13
DPPH(–) ; TEAC(–) ; RP(–) ; bC(–) DPPH(–)*3,6; RP(–)*6; HOSC(–)*6
EA W E (95%)
98.35 (GA) 60.88 (GA) 9.68 (PC)
DPPH(–)*7; RP(–)*7; HOSC(+)*7 DPPH(–)*7; RP(–)*7; HOSC(–)*7 DPPH(+)*1; HOSC(+)*1; RP(+)*1; TA(+)*1; LPPO(+)*1
Kaur et al., 2008
E (96%)
17.74
DPPH(–)*1,8; bC(–)*1,8
Moure et al., 2001
W
17.36
EA E W W
1.09 (GA) 1.94 (GA)
DPPH(–)*1,8; bC(–)*1,8 DPPH(–)*1,(+)*6; FIC(–)*3; RP(–)*1 DPPH(–)*1,(+)*6; FIC(–)*3; RP(–)*1; FTC(+)*1 DPPH(–)*1,6; RP(–)*1 DPPH(–)*2,3; RP(–)*2,3; TEAC(–)*2
Dimocarpus longan Lour. (peel) Gevuina avellana (seeds) Kappaphycus alvarezii
Litchi chinenesis (flowers)
9.92
*2
*2
*2,3
*5
*2,3
*5
*4
Shyu et al., 2009
*2
Liu and Yao, 2007 Marimuthu et al., 2008
*5
Yin et al., 2008
Pan et al., 2008b
Kumar et al., 2008
Liu et al., 2009
528 Separation, extraction and concentration processes
Table 18.7 Examples of extraction of plant food antioxidants using conventional solvent extraction. Symbols (+), (–) and (=) indicate higher/ lower/equal activity compared with the reference standard compound
W
1.60
DPPH(–)*1,8; bC(–)*1,8
W
30.96
FTC(+)*6; bX(+)*7(–)*8; DPPH(–)*7,8; SRS(–)*7; HOSC (–)*7,8; HPS(–)*7,8; RP(–)*7,8; FIC(–)*4
E
34.46
FTC(+)*6; bC(+)*7(–)*8; DPPH(–)*7,8; HOSC(+)*7(–)*8; HPS(–)*7,8; RP(–)*7,8; FIC(–)*4
EA
10.61
E
37 (GA)
FTC(+)*6, bC(+)*7(–)*8, DPPH(–)*7,8; HOSC(+)*7(–) *8; HPS(–)*7,8; RP(–)*7,8; FIC(–)*4 DPPH(–)*2,3,5,7,12(+)*1; SRS(+)*2,12; HOSC(–)*1,3,5 (+)*2,12
E
34 (GA)
DPPH(–)*1,2,3,5,7,12; SRS(–)*2,12; HOSC(–)*1,2,3,5,12
E
27 (GA)
DPPH(–)*1,3,7(+)*2,5,12; SRS(–)*2,12; HOSC(–)
W
–
FTC(+)*6; RP(+)*6; SRS(+)*1,6,8; DPPH(–)*5(+)*8; FIC(+)*1,6,8; HPS(–)*1,6,8
W E (50%)
8, 5.4 (GA) 12, 16.8 (GA)
Citrus sinensis
E EA
10, 10.4 (GA) 9.12, 6.43 (C)
Citrus grandis
EA
3.41, 9.72 (C)
Randia echinocarpa W (pulp) Rosa rubiginosa (seeds) E (96%) Smilax excelsa (leaves)
Syzygium cumini (kernel) Syzygium cumini (pulp) Syzygium cumini (seed coat) Urtica dioica (aerial parts)
9.46
*1,2,3,5,12
Ozsoy et al., 2008
Benherlal and Arumughan, 2007 Benherlal and Arumughan, 2007 Benherlal and Arumughan, 2007 Gülçin et al., 2004
Extraction of natural antioxidants from plant foods 529
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1.97
DPPH(–)*3,6, 8,9,10; TBARS(–)*6,8; SRS(–)*11; HOSC(–)*8 Ningappa et al., 2007 DPPH(+)*2,3,8,9,10; TBARS(+)*6,8; SRS(+)*11; HOSC(+)*8; RP(+)*8; FIC(–)*4 DPPH(–)*3,6, 8,9(+)*10; TBARS(+)*6,8; SRS(–)*11; HOSC(–)*8 DPPH(–)*3; ORAC(–)*12; ABTS(–)*3; RP(–)*2; PM(+)*13 Jayaprakasha et al., 2008 *3 *12 *3 *2 *13 Jayaprakasha et al., DPPH(–) ; ORAC(–) ; ABTS(–) ; RP(–) ; PM(+) 2008 *1 Santos–Cervantes bC(–) et al., 2007 *1,8 *1,8 Moure et al., 2001 DPPH(–) ; bC(–)
Murraya koenigii (leaves)
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*Total extractable polyphenols. E, ethanol; EA, ethyl acetate; W, water. C, catechin; GA, gallic acid; PC, pyrocatechol. *1 , butylhydroxytoluene (BHT); *2, (+)-catechin; *3, ascorbic acid; *4, ethylene diamine tetra-acetic acid (EDTA); *5, quercetin;*6, a-tocopherol; *7, gallic acid; *8, butylhydroxyanisol (BHA); *9, curcumin; *10, b-carotene;*11, superoxide dismutases (SOD); *12, trolox; *13, propyl gallate. bC, b-carotene bleaching method; DPPH, a,a-diphenyl-b-picrylhydrazyl radical scavenging assay; FIC, ferrous ion-chelating assay; FTC, antioxidant activity in the linoleic acid system with ferrothiocyanate reagent; HOSC, hydroxyl radical scavenging assay; HPS, hydrogen peroxide scavenging assay; LPPO, lipid peroxidation in peanut oil; ORAC, oxygen-radical absorbance capacity; PM, phosphomolybdenum method; RP, reducing power; SRS, superoxide radical scavenging assay; TA, total antioxidant capacity assay; TBARS, thiobarbituric acid-reacting substances; TEAC, Trolox equivalent antioxidant capacity assay.
530 Separation, extraction and concentration processes
Table 18.7 Continued
Extraction of natural antioxidants from plant foods 531 (enzyme-assisted aqueous extraction, some novel methods including subcritical water, pressurized liquid extraction, supercritical fluid extraction (SFE) with CO2, and ultrasound- and microwave-assisted processes) are considered in the following subsections. 18.5.2 Enzyme-aided extraction Enzyme processing is used as a pretreatment for conventional and alternative solvent extraction technologies, owing to the ability of enzymes to disrupt cell walls. Enzyme-aided extraction is considered more environmentally friendly than chemical treatments, and the further utilization of both residue and extract is possible (Mandalari et al., 2006). Enzyme treatments have been considered for a variety of materials, including algae, fruits, herbs and cereals. Edible and other abundant and underutilized seaweeds have been studied for extracting water-soluble compounds with antioxidant activity. Most antioxidant compounds are soluble in polar solvents, but the large amounts of highly viscous polysaccharides act as a physical barrier. Enzymes can break down the cell walls and storage materials to release both free compounds and polysaccharides attached to polyphenols (Heo et al., 2005). Representative studies of antioxidant extraction from algae biomass using hydrolytic enzymes are summarized in Table 18.8. Synergistic effects between different enzyme activities (including endopeptidases, carbohydrases and proteases) enhance the release of antioxidants (Siriwardhana et al., 2008; Heo et al., 2005). The variables affecting the extraction yield and activity include: type of raw material and pre-conditioning stages, temperature, type and concentration of solvent, and type and concentration of enzyme(s). This latter effect is shown in Fig. 18.2. Incubation results in increased extraction efficacy (Siriwardhana et al., 2008). Thermal degradation of high-MW polysaccharides enables higher extraction of target antioxidants. Some enzymatic hydrolysates are stable at 100 °C (Athukorala et al., 2006; Heo et al., 2005) and the DPPH radical scavenging activity increases with heating time (Siriwardhana et al., 2008). Heating followed by enzymatic hydrolysis enhances the access of endopeptidases and b-glucanase to laminarin and proteins. Increased antioxidant activity at alkaline pH was observed in Hizikia fusiformis, owing to hydrolysis, ion exchange, desulfation and prevention of the formation of the protein–polyphenol complex (Siriwardhana et al., 2008). A simple procedure to reduce the viscosity of porphyrans of (sulfated polysaccharide of Porphyra) various MW, consisting of treatments with hydroxyl radical, has been reported (Zhang et al., 2003). Algal enzymatic extracts show a variety of other biological activities. The enzymatic extract of Ecklonia cava inhibited proliferation of cancer cells strongly and selectively (Athukorala et al., 2006), and showed in vivo anti-inflammatory, inmunomodulatory and antiproliferative activity (Ahn et al., 2008). Highly sulfated fucose and galactose polysaccharides from
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Table 18.8 Examples of enzyme-aided extraction of plant food antioxidants
Algae Ecklonia cava Hizikia fusiformis © Woodhead Publishing Limited, 2010
Ishige okamurae Sargassum fullvelum Sargassum horneri Sargassum coreanum Sargassum thunbergii Scytosipon lomentaria Undaria pinnatifida
Terrestrial plants Buckwheat (Fagopyrum esculentum) Caroleo (Olea europaea)
Enzymes
Extraction yield (total phenolics)
Antioxidant activity
Reference
Viscozyme Alcalase 2.4L FG Alcalase and Ultraflo
1352 mg GAE g–1 1095 mg GAE g–1
DPPH; HO˙; O2˙; H2O2
Heo et al., 2005
DPPH; H2O2
Ultraflo L Alcalase 2.4L FG Ultraflo L Alcalase 2.4L FG Ultraflo L Alcalase 2.4L FG Viscozyme Protamex Viscozyme Alcalase 2.4L FG Ultraflo L Alcalase 2.4L FG Alcalase, Flavourzyme, Promozyme, Viscozyme, Celluclast
275 mg GAE g–1 420 mg GAE g–1 313 mg GAE g–1 366 mg GAE g–1 384 mg GAE g–1 533 mg GAE g–1 1123 mg GAE g–1 996 mg GAE g–1 386 mg GAE g–1 416 mg GAE g–1 149 mg GAE g–1 207 mg GAE g–1
DPPH; HO˙; O2–; H2O2
Siriwardhana et al., 2008 Heo et al., 2005
DPPH; HO˙; O2–; H2O2
Heo et al., 2005
DPPH; HO˙; O2–; H2O2
Heo et al., 2005
DPPH; HO˙; O2–; H2O2
Heo et al., 2005
DPPH; HO˙; O2–; H2O2
Heo et al., 2005
DPPH; HO˙; O2–; H2O2
Heo et al., 2005
DPPH; HO–; O2–
Je et al., 2009
Alcalase
0.371 mg R g–1
DPPH; reducing power; inhibition of LA peroxidation
Tang et al., 2009
Bioliva
234.1 mg CA kg–1
Oxidative stability; peroxide value Ranalli et al., 2003b
532 Separation, extraction and concentration processes
Vegetal material (Latin name)
Leccino (Olea europaea) Leccino (Olea europaea) Lemon peel (Citrus limon cv. Meyer) Lemon peel (Citrus limon cv. Yen Ben) Mandarin peel (Citrus reticulara cv. Ellendale)
Bioliva
362.2 mg CA kg–1
Oxidative stability; peroxide value Ranalli et al., 2003a
Rapidase adex D
128 mg CA kg–1
Oxidative stability; peroxide value Ranalli et al., 2003a
Cytolase
170 mg CA kg–1
Oxidative stability; peroxide value Ranalli et al., 1999
Cytolase
79 mg CA kg–1
Oxidative stability; peroxide value Ranalli et al., 1999
Rapidase adex D
92 mg CA kg–1
Oxidative stability; peroxide value Ranalli et al., 2003
Cellulase
220 mg F g–1
Huang et al., 2009
Pectinases and cellulases
89 mg CA g–1
HO˙; O2–; inhibition of lipid peroxidation; reducing power DPPH; oxidative stability
Cellulase MX Cellulase CL Kleerase AFP Bioliva
1.52 mg GAE g–1 1.30 mg GAE g–1 1.25 mg GAE g–1 121.2 mg CA kg–1
FRAP
Li et al., 2006
Cytolase
121 mg CA kg–1
Oxidative stability
Ranalli et al., 1999
Cellulase MX Cellulase CL Kleerase AFP Cellulase MX Cellulase CL Kleerase AFP Cellulase MX Cellulase CL Kleerase AFP
0.4429 mg GAE g–1 0.4895 mg GAE g–1 0.4375 mg GAE g–1 1.131 mg GAE g–1 0.99 mg GAE g–1 1.108 mg GAE g–1 1.39 mg GAE g–1 1.15 mg GAE g–1 1.19 mg GAE g–1
FRAP
Li et al., 2006
FRAP
Li et al., 2006
FRAP
Li et al., 2006
Ramadan et al., 2008
Oxidative stability; peroxide value Ranalli et al., 2003
Extraction of natural antioxidants from plant foods 533
© Woodhead Publishing Limited, 2010
Coratina (Olea europaea) Coratina (Olea europaea) Coratina (Olea europaea) Dritta (Olea europaea) Dritta (Olea europaea) Eucommia leaf (Folium eucommiae) Goldenberry (Physalis peruviana) Grapefruit peel (Citrus x paradisi)
© Woodhead Publishing Limited, 2010
Vegetal material (Latin name)
Enzymes
Extraction yield (total phenolics)
Antioxidant activity
Moringa concanensis seeds
1.283 mg T kg–1 1.236 mg T kg–1 1.176 mg T kg–1
Oxidative stability; peroxide value; Latif and Anwar, conjugated diene/triene 2009
Sunflower (Helianthus annuus)
Kemzyme Natuzyme Feedzyme Protex 7L Kemzyme Alcalase 2.4L Viscozyme L Natuzyme
Sweet orange peel (Citrus sinensis cv. Navel)
Cellulase MX Cellulase CL Kleerase AFP
0.80 mg GAE g–1 0.79 mg GAE g–1 0.90 mg GAE g–1
FRAP
Li et al., 2006
Macer8 Fj (M) Grindamyl pectinase (G) (M) + (G) Olivex and Celluclast
358 mg GAE L–1 237 mg GAE L–1 383 mg GAE L–1 5.2 g C kg–1
Conjugated diene
Landbo and Meyer, 2001
DPPH
Soto et al., 2008
Grindamyl pectinase (G) Celluclast (C) (G) + (C) Ultrazym 100G and Cellubrix Ferulic acid esterase
3072 mg GAE L–1 2773 mg GAE L–1 3017 mg GAE L–1 0.25 g C g–1
Conjugated diene
Meyer et al., 1998
DPPH
Collao et al., 2007
6937 mg SAE g–1
DPPH; conjugated diene; hydroperoxide; hexanal; liposome model system
Vuorela et al., 2004
Residues Blackcurrant pomace (Ribes nigrum) Borage seeds (Borago officinalis) Grape seeds pomace (Vitis vinifera) Evening primrose seeds (Oenothera biennis) Rapeseed (Brassica rapa)
13 14 13 15 13
mg mg mg mg mg
GAE GAE GAE GAE GAE
g–1 g–1 g–1 g–1 g–1
Reference
DPPH; inhibition of LA; oxidation Latif and Anwar, 2009
C, catechin; CA, caffeic acid; F, flavonoid; GAE, gallic acid equivalents; R, rutin; SAE, sinapic acid equivalents; T, tocopherol.
534 Separation, extraction and concentration processes
Table 18.8 Continued
Extraction of natural antioxidants from plant foods 535 Extractable compounds (mg/100 g)
4.5
1
4.0 3.5
2
3
4
3.0 2.5 5
2.0 1.5
7
1.0
9
0.5 0
6
8
10 0
2
4 6 Enzyme concentration (%)
8
10
Fig. 18.2 Effect of enzyme concentration on the yield of extractable compounds. 1: Bilberry (Biopectinase CCN) (Puupponen-Piniä et al., 2008). 2: Bilberry (Pectinex BE 3L) (Puupponen-Piniä et al., 2008). 3: Blackcurrant (Biopectinase CCN) (PuupponenPiniä et al., 2008). 4: Blackcurrant (Pectinex BE 3L) (Puupponen-Piniä et al., 2008). 5: Pigeonpea (Fu et al., 2008); lutein (mg/100 g) ¥ 10–1; enzyme concentration ¥ 10–2. 6: Grape pomace (Maier et al., 2008); enzyme concentration ¥ 10–1; 7: Apple pomace (Pinelo et al., 2008). 8: Pigeonpea (Fu et al., 2008); apigenin ¥ 10–1; enzyme concentration ¥ 10–2. 9: Meyer lemon (Li et al., 2006). 10: Yen Ben lemon (Li et al., 2006).
Ecklonia cava showed antiproliferative action (Athukorala et al., 2006). The phlorotannin derivatives of this alga exhibited inhibitory effect on human immunodeficiency virus type 1 reverse transcriptase and protease (Ahn et al., 2004). Enzymatic extracts of Stellaria dichotoma inhibited the in vitro hydroxyl radical-induced DNA damage (Lim et al., 2008). A peptide with ACE-inhibitory activity was isolated from the pepsin hydrolyzate of Chlorella vulgaris waste (Sheih et al., 2009), whereas some glucans (hexaose, schizophyllan, and laminarin) showed inhibition against AAPH (Tsiapali et al., 2001). The major antioxidants of fruits and juices are phenolic compounds. Phenols may be found in cell walls, bound to polysaccharides by hydrophobic interactions and hydrogen bonds, linked to the protein matrix of vacuolar inclusions, confined inside the cellular vacuoles, and near or associated with the cell nucleus (Pinelo et al., 2006). Plant cell walls are a complex network of cellulose, hemicelluloses (mainly of xyloglucans, mannans, xylans, and arabinogalactans), pectins and lignin (a hydrophobic polymer derived from p-coumaryl, coniferyl, and sinapyl alcohols). Degradation of the cell-wall polysaccharide structure is a fundamental step in the release of phenols linked to the cell wall or associated with cell vacuoles. Ferulic and p-coumaric acids, the major lignin monomers that link hemicellulose sugars and lignin, are potent in vitro antioxidants (Meyer et al., 1998; Nardini et al., 1995). Enzymatic degradation of cell-wall polysaccharides is assumed to increase © Woodhead Publishing Limited, 2010
536 Separation, extraction and concentration processes the overall substrate porosity, facilitating solvent penetration and extraction efficiency, but lignin and tannins may block the enzyme action. The use of enzymes that degrade plant cell walls to enhance the extraction of phenolic antioxidant compounds has been reported for food products and by-products. Although no interactions between proteases and pectinases was confirmed for blackcurrant juice production, degradation of the protein network enhanced both substrate porosity and phenol recovery (Landbo and Meyer, 2001). In grape skins, cellulases, hemicellulases, pectinases, and other hydrolytic enzymes can be used (Pinelo et al., 2006). Despite the high cellulose content of the grape cell wall, cellulolytic activity did not influence the extraction yield of phenols from pomace (Meyer et al., 1998), but pectinases, cellulases and proteases increased the extraction yields from apple skin (Pinelo et al., 2008). Preferential enzymecatalyzed solubilization of pectin and cellulose from bergamot peel caused the selective release of glucose and galacturonic acid (Mandalari et al., 2006). Glucanases and pectinases were used for processing citrus peels (Li et al., 2006). Neutral protease, papain and alkaline protease have been assayed on the fruit of Physalis alkekengi (Ge et al., 2009), whereas culture broths of Aspergillus niger enriched in cinnamoyl esterases were used to release hydroxytyrosol from olive oil by-products (Bouzid et al., 2005). The combined effect of the major operational variables (including particle size, time, temperature, and enzyme to substrate ratio) is evaluated using experimental designs (Meyer et al., 1998). The particle size is relevant for degrading plant cell wall carbohydrates, as well as for solvent penetration and solute recovery. A favorable effect of the reduction of particle size on the extraction of phenolics was reported for spent coffee grounds (Pinelo et al., 2007), blackcurrant pomace (Landbo and Meyer, 2001), apple skin (Pinelo et al., 2008), and citrus peels (Li et al., 2006). Reduced particle size favored the degradation of polysaccharides by pectinase without affecting the recovery of phenols (Meyer et al., 1998). Although a beneficial effect of high enzyme concentration (5%) was reported for citrus peel, the optimal value for this material was 1–3% (Li et al., 2006). Increased enzyme dosage enabled higher juice yields and phenol concentrations from blackcurrant mash. However, the non-enzyme-treated juices exhibited higher antioxidant activity than the enzyme-treated ones, presumably because of differences in their phenolic profiles (Bagger-Jørgensen and Meyer, 2004). The effects of enzymatic treatments were different for various citrus fruits (lemon, mandarin, grapefruit and orange) (Li et al., 2006), and various effects of temperature, particle size reduction and ethanol concentration were reported for the extraction of individual compounds from apple skins (Pinelo et al., 2008). Optimal temperatures of 50–60 °C were reported for citrus peel (Li et al., 2006), apple skins (Pinelo et al., 2008), and elderberry (Landbo et al., 2007). Prolonged incubation time (24 h) with pectinolytic and cellulolytic enzymes was necessary to solubilize carbohydrates and flavonoids from bergamot peel (Mandalari et al., 2006), but a negative effect of long reaction
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Extraction of natural antioxidants from plant foods 537 times was observed for grape owing to degradation of phenolic compounds (Meyer et al., 1998). Enzyme-assisted aqueous of citrus peels required an incubation period of 3 h, compared with 6 h for water extraction without enzymes (Li et al., 2006). The amount of quercetin in enzyme-treated juices increased with incubation time, but the quercetin content did not involve a higher antioxidant activity (Sun et al., 2005). The yields of soluble carbohydrates were higher in enzyme-treated samples than in untreated ones. A positive linear correlation between extent of cell wall degradation and phenol extraction was observed for grape pomace (Meyer et al., 1998), elderberry (Bagger-Jørgensen and Meyer, 2004), and blackcurrants (Bagger-Jørgensen and Meyer, 2004; Landbo and Meyer, 2001), but no significant differences were found in carbohydrate degradation, distribution of phenolics and release of anthocyanins and total phenols (Landbo et al., 2007). Similarly, pectinolysis affected both juice yield and contents of phenolic acids, flavonoids and anthocyanins (Kaack et al., 2008). Even though the amount of phenols extracted from the pomace is enhanced, specific compounds may decrease: for example, enzyme-mediated anthocyanin degradation may occur owing to enzymatic hydrolysis of glycosylated anthocyanins by polyphenol oxidases. Pectinase preparations produced by Aspergillus niger, containing b-glucosidase, b-galactosidase, and a-l-arabinosidase, affect anthocyanin pigments and color quality (Landbo and Meyer, 2001). Despite b-galactosidase and b-glucosidase causing color degradation, benefits for extraction of potent antioxidants from grape pomace were reported (Meyer et al., 1998). Enzymatic release of esterified hydroxycinnamic acids can be achieved by synergistic action of cell-wall-degrading enzymes (esterases and xylanases). Commercial pectinases facilitate the release of ferulic acid and other phenolic acids from ground rye grain (Andreasen et al., 1999). A cold pressing process was used for defatting borago before extracting phenolics with conventional solvents. Enzymatic pretreatment of seeds enhanced the extraction of phenolic compounds with DPPH radical scavenging capacity from the meal; the tocopherol content of the resulting oil was maintained by cold pressing (Soto et al., 2008). Treatments with cell-wall-degrading enzymes were proposed for releasing ferulic acid from agro-industrial byproducts such as sugar beet pulp or maize bran (Bouzid et al., 2005). Other less studied materials include microbial sources. Saccharomyces cerevisiae b-glucan, isolated by hot water and enzyme treatment, exhibited antioxidant activity (Jaehrig et al., 2007). Protease treatments can be included within the enzyme-assisted extraction methods. Relatively short bioactive peptides (2–9 amino acids) can be obtained from sources such as wheat, corn, soybean and mushrooms (Dziuba et al., 2004). Commercial proteases have been used to process a number of substrates (Tang et al., 2009; Wang et al., 2007). Hydrolysis increased ACE inhibition and the radical-scavenging activity of protein hydrolyzates from potato isolates and by-products, possibly related to peptides and /or
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538 Separation, extraction and concentration processes free amino acids liberated during digestion (Pihlanto et al., 2008). Rapeseed protein hydrolyzates present DPPH radical-scavenging activity, provide protection against the peroxidation of lipids and/or fatty acids (Amarowicz and Shahidi, 1997; Peña-Ramos and Xiong, 2002; Sakanaka et al., 2004) and show reducing power and Fe2+ chelating ability (Moure et al., 2006; Tang et al., 2009). Esterified phenolic acids (p-coumaric acid, ferulic acid) present in sugarcane bagasse were released by alkaline hydrolysis (Ou et al., 2009). Ferulic acid (FA), a scavenger of free radicals approved in certain countries as a food additive to prevent lipid peroxidation (Srinivasan et al., 2007), can be extracted using enzymatic, alkaline or acidic extractions. 18.5.3 Pressurized-liquid extraction Pressurized-liquid extraction (PLE) is an efficient, innovative and environmentally clean technique, which is performed at high temperature and pressure to maintain the solvent in the liquid state. Compared with traditional extraction techniques, it is faster and requires lower amounts of solvent. Alternatively, it is known as accelerated solvent extraction (ASE), or high-pressure liquid extraction (HPE) (Adil et al., 2008). The structure, activity and properties of the extracts are usually unaffected in an oxygenfree and light-free environment, and micro-organisms and enzymes may be inactivated (Qadir et al., 2009). The influence of HPE on the antioxidant capacity of food products was recently reviewed by Oey et al. (2008). PLE was proposed for the extraction of a variety of natural products and wastes, including the production of anthocyanins from grape skins (Ju and Howard 2003), flavonoids from spinach (Howard and Pandjaitan, 2008) and processing of berry substrates (King et al., 2003). The use of non-toxic solvents (water and ethanol) offers environmental advantages. Ethanol–water mixtures gave high extraction yields from Phormidium species (Rodríguez-Meizoso et al., 2008), sour cherry pomace (Adil et al., 2008), and dried spinach (Howard and Pandjaitan, 2008); whereas methanol was used for extracting catechin and epicatechin from tea leaves and grape seeds (García-Marino et al., 2006; Piñeiro et al., 2004), and flavonoids from apple extracts (Alonso-Salces et al., 2001). The compounds recovered and their extraction efficiency differed with the solvent, providing different properties to the extracts, including color and odor compounds (Howard and Pandjaitan, 2008). The amount of carotenoids extracted from Dunaliella salina with hexane was more than seven times higher than the value extracted with ethanol, but the antioxidant activity was only twice as high owing to structural differences (Herrero et al., 2006). Combinations of pressure, temperature, solid/solvent ratio and extraction time can be addressed by experimental designs (Adil et al., 2008; Herrero et al., 2006). Temperature, vapor pressure of the most volatile compounds and improved mass transfer have been reported as the major variables involved in
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Extraction of natural antioxidants from plant foods 539 PLE regardless the solvent used (Rodríguez-Meizoso et al., 2008). Temperature was the most influential factor in the extraction of antioxidant compounds from Dunaliella salina (Herrero et al., 2006). The optimal temperature for water and ethanolic extraction of Spinacia oleracea flavonoids was 50–130 °C for water, and 50–150 °C for ethanol (Howard and Pandjaitan, 2008), compared with 100–130 °C for PLE processing of grape by-products (Ju and Howard 2003; Piñeiro et al., 2004) and berries (King et al., 2003). The optimal temperature depends on the solvent considered. Optimal antioxidant activity of Phormidium species in hexane extracts was obtained operating at 50 °C, whereas in ethanol and water extracts it was observed at 150 and 200 °C, respectively (Rodríguez-Meizoso et al., 2008). Thermal degradation of apple polyphenols was observed at extraction temperatures higher than 60 °C in methanol, whereas activity decreased at 110–190 °C in water and ethanol (Alonso-Salces et al., 2001). Thermal degradation of flavonoids was reported for extraction temperatures >130 °C (Howard and Pandjaitan, 2008). The reduction in antioxidant activity observed at prolonged extraction times during PLE of Dunaliella salina with hexane and ethanol suggested some kind of carotenoid degradation (Herrero et al., 2006). HPE extraction can give a moderately high extraction efficiency of bioactive compounds in a short time, but treatments for prolonged periods could result in degradation (Qadir et al., 2009). The development of brown, highly polar, odoriferous compounds with high antioxidant capacity at high extraction temperatures suggested the participation of Maillard reactions in ethanolic (Howard and Pandjaitan, 2008) and in water extracts (Rodríguez-Meizoso et al., 2008). Extracts of propolis obtained by high hydrostatic pressure extraction in 1 min presented similar properties to the ethanolic extracts obtained by leaching at room temperature for a few days (Jun, 2006). No degradation of compounds present in ethanolic extracts occurred after exposure to light and air at room temperature for 2 days (Herrero et al., 2006). Other reported biological activities include cytotoxic effects on the growth of human promyelocytic leukemia cells (HL-60) of sweet potato extracts (Rabah et al., 2005), toxicity on preformed monolayers of Vero cells (African green monkey kidney cell line) of Phormidium extracts from PLE (Rodríguez-Meizoso et al., 2008). Antimicrobial activity of Phormidium sp. extracts has been ascribed to terpenes (i.e., b-ionone, neophytadiene) and fatty acids (i.e., palmitoleic and linoleic acids) (Rodríguez-Meizoso et al., 2008). Subcritical-water extraction (SWE) is an environmentally friendly technology for the selective extraction of bioactive compounds from plant materials with advantages derived from its simplicity, reduced extraction time, high quality of the extract, low cost of the extracting solvent and possibility of scaling-up (Rodríguez-Meizoso et al., 2006). Extract with water under pressure (between 100 °C and the critical temperature of 374 °C) is also known as superheated water extraction, subcritical water extraction, high-temperature water extraction, pressurized hot water extraction, or hot
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540 Separation, extraction and concentration processes water extraction. Most studies have been performed in batch mode on a small scale, but continuous flow equipment has also been described (Budrat and Shotipruk, 2008). The feedstocks studied include fruits or vegetables and their processing wastes, seeds, herbs and algae (Table 18.9). SWE was more efficient than conventional solvents for extracting flavanols from grape seeds (García-Marino et al., 2006) or antioxidant phenolics from bitter melon (Budrat and Shotipruk, 2008), and was sucessfully employed to concentrate and isolate antioxidant compounds from oregano leaves (Rodríguez-Meizoso et al., 2006). Temperature has a marked effect on yield and selectivity. The dielectric constant of water decreases with temperature, enabling the extraction of compounds having different polarities at different temperatures. High temperature also enhances diffusivity, facilitating the transport of solutes from the solid matrix, and a compromise to avoid thermal degradation must be reached. The release of hydroxycinnamates from cell walls is favored by temperature, but anthocyanins undergo degradation reactions. Lignin decomposes in sub- and supercritical water to give phenol, which are degraded at high temperatures (Budrat and Shotipruk, 2008; Garrote et al., 2003). Sequential extraction with stepwise pressure increase was proposed for processing black tea leaves (Chambers et al., 1984) as well as to recover quercetin glycosides from onion waste (Turner et al., 2006), and catechins and proanthocyanidins from winery by-products (García-Marino et al., 2006). The major variable is temperature, because at 150 °C or above, substrates are hydrolyzed, and the solvent polarity decreases. Optimization of operational conditions has to be carried out for each material. The yield of total phenolics from sweet potato and bitter melon increased with temperature (Budrat and Shotipruk, 2008; Rabah et al., 2005). Although higher temperatures resulted in higher amounts of extracted phenolics, their antioxidant activity was higher operating at lower temperatures (Budrat and Shotipruk, 2008). Carnosic acid, the most potent antioxidant in rosemary, was preferentially extracted at 200 °C (Rodríguez-Meizoso et al., 2006). Extraction yields of oregano leaves increased with temperature: at the lowest temperature considered, the more polar compounds (flavanones and dihydroflavonol structures such as dihydroquercetin, eriodictyol and dihydrokaempferol) were preferentially extracted, whereas less polar compounds could also be extracted at 200 °C (Rodríguez-Meizoso et al., 2006). The phenolic content was a maximum at 25 °C and a minimum at 200 °C, but the highest antioxidant activities were achieved at 150–200 °C. Boiling ground coffee beans in water at 110 °C under elevated pressure gave an efficient extraction of antioxidants (Budryn and Nebesny, 2008). The total phenolic contents and the antioxidant activities of HPE extracts from longan fruit were higher than those obtained with conventional solvent extract (Prasad et al., 2009), whereas enhanced extraction of onion flavonols was observed at 100–400 MPa (Roldán-Marín et al., 2009). The conditions of SWE influence the degree of polymerization and structure
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Table 18.9 Examples of extraction of plant food antioxidants using high pressure systems Raw material
Operational conditions
Elderberry leaves (Sambucus nigra L.) Elderberry berries (Sambucus nigra L.) Elderberry flowers (Sambucus nigra L.) Phormidium species
Red grape skin (Vitis vinifera)
Antioxidant activity Reference
0.32 mg GAE g–1
DPPH
Bonoli et al., 2004
Fiber: 14.99 g GAE kg–1 Flour: 20.93 g GAE kg–1
DPPH
Papagiannopoulos et al., 2004
28.91 mg GAE g–1
DPPH ABTS FRAP DPPH bC bleaching
Güçlü-Üstündağ and Mazza, 2009
DPPH bC bleaching DPPH bC bleaching TEAC Antimicrobial activity
Dawidowicz et al., 2006 Dawidowicz et al., 2006 Rodríguez-Meizoso et al., 2008
ORAC
Ju and Howard, 2003
1 g; ECV(22); 17 (g F/100 g ¥ 10–2) SE: E:W (80:20); -; 10 min; 60 bar; 200 ºC; PLE: FV( 60% ECV); P(120) 20.18 (g F/100 g ¥ 10–2) 214.2 (g F/100 g ¥ 10–2) 1 g sample H: 20 min; 200 ºC; E: 20 min; 200 ºC; W: 20 min; 200 ºC; 0.2 g + 30 g sea sand; ECV(22); SE: acidified W; 3 cycles; 5 min; 10.1 MPa; 50 ºC PLE; FV (50% ECV); P(90)
6.54% 40.91% 22.81% 179.9 mg GAE g–1
Dawidowicz et al., 2006
Extraction of natural antioxidants from plant foods 541
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Pressurized liquid extraction (PLE) Barley flour 2 g + 4 g Hydromatrix; ECV(33); (Hordeum vulgare L.) SE: E:W (4:1); 2 cycles; 5 min, 20 MPa 60 °C PLE: FV (60% ECV); P(60) Carob fruits 1 g + 2 g diatomaceous earth; (Ceratonia siliqua L.) SE: A:W (1:1); 2 cycles; 5 min; -; 60 °C; PLE: 50% FV; Cow Cockle seeds 2 g; (Saponaria vaccaria) W; 3 h; 125–175 °C; F(2)
Solubles yield or TEP*
Raw material
Operational conditions
Spinach (Spinacia oleracea)
0.5 g + 30 g Ottawa sand 42 mmol GAE kg–1 SE: E:W (7:3); 3 cycles; 5 min; 13.8 MPa; -; PLE: R (70% ECV) 2.5 g; ECV(11) PE; 9 min; 170 ºC; 77.0 mg extract H; 15 min; 170 ºC; 95.5 mg extract E; 9 min; 170 ºC; 270.1 mg extract W; 15 min; 170 ºC; 290.0 mg extract 2.5 g; ECV(11) PE; 9 min; 170 ºC; 2.94% H; 9 min; 170 ºC; 4.28% E; 9 min; 170 ºC; 19.70%
Spirulina plantensis © Woodhead Publishing Limited, 2010
Spirulina plantensis
Solubles yield or TEP*
Antioxidant activity Reference ORAC
DPPH
DPPH
Howard and Pandjaitan, 2008 Herrero et al., 2004
Jaime et al., 2005
High pressure extraction (HPE) Korean barberry bark (Berberis koreana)
1 g mL–1; 15 min; 500 MPa; RT
50% E; 0.02 g mL–1; 30 min; 500 MPa; 50 ºC E; 0.06–0.07 g mL–1; 25 min; 176–193 MPa; 60 °C High pressurized fluid extraction (HPFE) Brazilian propolis RSL 1:10; 20% E; 40 g surfactant; 50 psig; 393 K Longan fruit pericarp (Arillus longan) Sour cherry pomace
11.04% (317.35 mg GAE g–1)
DPPH Inhibition of xanthine oxidase
Qadir et al., 2009
30% (23 mg GAE g–1)
Prasad et al., 2009
3.80 mg GAE g–1
DPPH O2– DPPH
Adil et al., 2008
25.2%
DPPH
Chen et al., 2007
542 Separation, extraction and concentration processes
Table 18.9 Continued
Taiwan Red Pine (Pinus taiwanensis)
70% E; 180 min; 689 kPa; 343 K; F (10)
26.3 mg g–1
Taiwan White Pine (Pinus morrisonicola)
70% E; 180 min; 689 kPa; 343 K; F (10)
39.8 mg g–1
Propolis crude
RSL 1:35, E 75%; 1 min; 500 MPa; RT
Accelerated solvent extraction (ASE) Spirulina plantesis 2.5 g; ECV (11); 15 min; -; 170 °C
–
DPPH
Rieger et al., 2008
–
DPPH
Rieger et al., 2008
317 mmol GAE g–1 dm
ABTS
Corrales et al., 2008
–
DPPH
Rieger et al., 2008
TPA: 0.20, TF: 7.10, TOD: DPPH ORAC 63.5 mg GAE g–1
Súarez et al., 2009
290 mg GAE g–1
bC bleaching DPPH
Xi, 2006
6.43 g GAE g–1
bC bleaching DPPH
Xi and Shouqin, 2007
H: 4.3%; PE: 4.01% E: 17.14%; W: 10.12%
DPPH
Herrero et al., 2005
Extraction of natural antioxidants from plant foods 543
© Woodhead Publishing Limited, 2010
High hydrostatic pressure (HPP) Bilberry ECV(11); (Vaccinium myrtillus) SE: 3 cycle, 7 min, 60 bar, 80 ºC; PLE: FV (100%), P (100) Elderberry ECV (11); (Sambucus nigra) SE: 3 cycle; 5 min; 68.9 bar; 60 °C; PLE: FV (100%), P (60) Grape skins RSL 1:4.5; E:W (1:1); 1 h; 600 MPa; (Dornfelder V. vinifera) 70° C; Heather ECV (11); (Calluna vulgaris) SE: 3 cycle, 5 min; 60 bar; 60 ºC PLE: FV (100%), P (60) Olive cake 30 g OC; ECV (100) E:W (80:20), FV (60% ECV), 2 cycles; 5 min Propolis extract RSL 1:35, E 75%; 1 min; 500 MPa; RT
DPPH, Fe Lin et al., 2009 chelating, reducing ability, O2–, NO DPPH, Fe Lin et al., 2009 chelating, reducing ability, O2– NO
Raw material
Operational conditions
© Woodhead Publishing Limited, 2010
Solubles yield or TEP*
Antioxidant activity Reference
Hot water extraction (HWE) Canola meal RSL 1:16.7; W; 30 min; -; 80 °C (Brassica napus)
0.15 g g–1; 7.83 g SA g–1
Citrus peels
RSL 1:20; bddW; 30 min; -; 100 ºC
~32.5 mg GAE g–1 dw
Formosa koa leaves (Acacia confusa)
bdd H2O; 4 h; -; -
7.3 g
Green tea
RSL 1:100; dW; 5–30 min; -; 80 ºC
Herbal tea (Anagallis arvensis) Japanese horse chestnut (Aesculus turbinata ) Kaffir lime fruit peel (Citrus hystrix) Lettuce extract (Lactuca sativa) Mallotus japonicus leaves
RSL 1:166.7; bW; 5 min;- ; -
TP: 2050 mg GAE L–1 TF: 1700 mg GAE L–1 1.47 mmol Trolox g–1
RSL 1:100; bW; 2 h
1560 mg (280 mg EE)
DPPH TEAC Reducing power bC bleaching DPPH FRAP DPPH NBT Reducing power FRAP ABTS CUPRAC ABTS DPPH
0.2 g; ECV (10.8); 15 min; 200 ºC
23.7 mg GAE mL–1
DPPH
RSL 1:2; W; 80 ºC; 10 min
–
RSL 1:50 mL; bdW; 20 min
7.5 mg GAE mL–1
Morinda citrifolia roots
0.5g, ECV(10); 4 MPa; 2 h; 170 °C F (5)
92.55%
Peroxidation of liposomes DPPH O2– HO· DPPH
Hassas-Roudsari et al., 2009 Xu et al., 2008 Tung et al., 2009 Rusak et al., 2008 Apak et al., 2006 Ogawa et al., 2008 Khuwijitjaru et al., 2008 Altunkaya et al., 2008 Tabata et al., 2009 Pongnaravane et al., 2006
544 Separation, extraction and concentration processes
Table 18.9 Continued
Samor thai fruits (Terminalia chebula Retz.) Sasa palmata leaf
15.96 mg TP g–1 dw
ABTS
Rangsriwong et al., 2009
SE: 50 g; steam; 0.5–20 min; 180–260 °C (1.0–4.9 MPa) HWE: RSL 1:100; dW; 2 h; 98 °C RSL 1:100; dW; 2 h; 98 ºC
128.24 mg GAE g–1
DPPH
Kurosumi et al., 2007
12.08 mg GAE g–1
DPPH
Kurosumi et al., 2007 Spigno and De Faveri, 2009 Rusak et al., 2008
Tea
RSL 1:100; bW; 30–210 s
White tea
RSL 1:100; dW; 5–30 min; 80 ºC
Hot pressurized water extraction (HPWE) Boldo RSL 1:10; W; 3 h; 110 ºC (Peumus boldus M.) Sage 0.15 g + sea sand; 60 min; 100 kg cm–2; 100 ºC; F(1) (Salvia officinalis) Yam (Dioscorea alata)
LSR: 20 kg kg–1; ECV(1000); W; 3 h; 1.34 MPa; 120 ºC; F(10)
Subcritical water extraction (SWE) Bitter melon 1 g; ECV (10); (Momordica charantia) 3 mL min–1; 60 min, 200 °C, 10 MPa Canola meal 30 min; 6.89 MPa; 160 °C; F(1)
ABTS TP: 1300 mg GAE L–1 TF: 900 mg GAE L–1
FRAP ABTS
51.4% 5.2 mg B kg–1
ABTS DPPH
Del Valle et al., 2005 Ollanketo et al., 2002
2.143 g kg–1
DPPH
Chen et al., 2004
53 mg GAE g–1
ABTS
0.45 g g–1 15.42 g SA g–1
TEAC bC bleaching DPPH Reducing power
Budrat and Shotipruk, 2008 Hassas-Roudsari et al., 2009
Extraction of natural antioxidants from plant foods 545
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RSL 1:150; 100 °C; 120 min
© Woodhead Publishing Limited, 2010
Raw material
Operational conditions
Solubles yield or TEP*
Antioxidant activity Reference
Eucalyptus leaves (Eucalyptus globulus) Licorice roots (Glycyrrhiza glabra) Oregano (Origanum vulgare)
60 g; ECV (300) 200 mL W; 30 min; 4.91 MPa; 200 ºC 0.1 g; ECV (18.57) 10 mL W; 60 min; 0.002–5 MPa; 300 °C 0.75 g; ECV (11) SE:W; 15–30 min; 1500 psi; 25–200 °C; SWE: R (70% ECV) 0.40 g + 30 g sea sand; ECV (22) SE: 1 cycle; 40 s; – 110 ºC; SWE: R (70% ECV), P (90) 1.0 g; ECV (10) 37.5 min; 4 MPa; 180 °C; F (4) RSL 1:50; 10–60 min; 0.47–4.76 bar; 80–150 °C
4.8 mg g–1
Peroxynitrite
1521.3 ppm GAE 54% (0.149 g GAE g–1)
Radical scavenging Baek et al., 2008 Reducing power DPPH Rodriguez-Meizoso et al., 2006
52.3 mg TP g–1
ORAC
Ju and Howard, 2005
21.43 mg TP g–1
ABTS
0.73 mg TC g–1 dm 5.1 mg TF g–1 dm
ABTS
Rangsriwong et al., 2009 Prommuak et al., 2008
Red grape skin Samor thai fruits (Terminalia chebula) Thai silk waste *
Kulkarni et al., 2008
Total extractable polyphenols. B, boldine; bdW, boiled distilled water; bddW, boiled double distilled water; bW, boiled water; bC, b-carotene bleaching; dm, dry matter; dW, distilled water; E, ethanol; ECV, extraction cell volume; EE, epicatechin equivalent; F, flux (mL/min); FV, flush volume; GAE, gallic acid equivalent; H, hexane; LSR, liquid–solid ratio; P, purge (s); PE, petroleum ether; PLE, pressurized liquid extraction; pm, plant material; R, rising; RT, room temperature; SE, static extraction; TC, total carotenoids; TF, total flavonoids; TP, total phenolics; W, water.
546 Separation, extraction and concentration processes
Table 18.9 Continued
Extraction of natural antioxidants from plant foods 547 of molecules, as reported for procyanidin extraction (García-Marino et al., 2006). Enhanced recoveries of flavanol dimers and trimers, and gallic acid was obtained in greater quantities by a single extraction at 150 °C, in which gallic acid accounted for 61% of total phenolics. Higher temperatures favored the extraction of oligomeric fractions. When the galloylated moieties were situated in the terminal sub-unit, sequential extraction (50–100 °C) was preferable, whereas a single extraction at 150 °C was better for the rest of the galloylated derivatives (García-Marino et al., 2006). Treatment of lignocellulosic substrates with water or steam at 160–240 °C (autohydrolysis processing) results in depolymerization of hemicelluloses and in breakage of lignin-carbohydrate bonds, leading to solubilization of lignin fragments of low MW. The aqueous treatments of lignocellulosics for both the hydrolytic degradation of hemicelluloses and the solubilization of antioxidant compounds have been reviewed (Garrote et al., 2004; Meireles, 2009). Other related technologies, based on the utilization of chemicals (for example, solvents, oxygen or acids) added to water, were proposed to provide higher extraction yields. Steam treatments were efficient in releasing hydroxytyrosol and 3,4-dihydroxyphenylglycol (DHPG), from the semisolid waste of olive oil extraction systems (alperujo) (Fernández-Bolaños et al., 2002; Rodríguez et al., 2009). The amount of solubilized hydroxytyrosol increased with temperature and time, reaching 1.4−1.7 g/100 g of dry alperujo (FernándezBolaños et al., 2002). 18.5.4 Supercritical-fluid extraction Supercritical-fluid extraction (SFE), based on the utilization of a fluid under supercritical conditions, is suitable for the extraction and purification of natural compounds with bioactive and antioxidant properties. The interest in SFE is promoted by increasingly restrictive environmental, toxicological, and health regulations. Carbon dioxide shows favorable properties, including environmental safety, high availability at low cost, high purity, and suitability for extracting heat-labile, natural compounds with low volatility and polarity. The physicochemical properties of supercritical CO2 (high diffusivity, low viscosity and low surface tension compared with conventional solvents) facilitate mass transfer, enabling an environmentally friendly operation. The properties of the final product are improved by high selectivity, low processing temperatures, absence of light and oxygen, and absence of solvent in extracts. The extracts are regarded as natural, and do not require additional sterilization. The major disadvantages of supercritical CO2 are the high critical pressure (requiring expensive equipment), and the poor solvent power derived from its low polarity. The use of cosolvents enhances the solubility of some compounds and may increase the extraction selectivity, allowing operation at a lower pressure. Representative antioxidant compounds extracted by SFE are summarized
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548 Separation, extraction and concentration processes in Table 18.10. SFE successfully recovered phenolic compounds from liquid solutions (Gamse, 2004), and countercurrent contact has been employed for processing propolis (Wang et al., 2004), B. kaoi (Wang et al., 2005) and orange juice (Señoráns et al., 2001). Studies on the SFE of bioactive compounds with antioxidant activity are available (Del Valle et al., 2005; Díaz-Reinoso et al., 2006; Herrero et al., 2006; Meireles, 2003; Mukhopadhyay, 2000; Reverchon, 1997). The most influential variables in SFE of compounds with antioxidant properties from natural sources are summarized here. The drying technique (Mechanical and thermal conditioning) affects both the structure of the solid matrix and the thermal stability of the active compounds. Mass transfer is facilitated by a higher degree of grinding, but the performance of fixed beds during leaching could be limited by too fine particles. Pressure and temperature are the variables that determine the solubility equilibrium, an understanding of which is essential for process design. Results are mostly available for a single solute in the supercritical solvent, but the presence of other solutes may affect the equilibrium, and the natural products are a complex mixture of compounds. Studies dealing with the solubility of solid mixtures of phenolic compounds in supercritical CO2 have been reported (Del Valle et al., 1998; Díaz-Reinoso et al., 2006; Fornari et al., 2005; Lucien and Foster, 2000; Reverchon, 1997). Solvent density increases with pressure, but beyond a certain threshold, the increased solvent viscosity reduces the diffusion coefficients. When the target compounds are lipophilic, the operational conditions may be milder than those required for phenolic antioxidants. Highly labile compounds require mild temperatures (below 50 °C) to avoid alteration. A stepwise increase in the extraction pressure allows selective extraction of different compounds. The presence of a modifier increases the solvent density resulting in a higher interaction of solutes with the solvent, causing alterations and swelling of the vegetal matrix. A suitable cosolvent may improve the extraction yield and selectivity. However, cosolvents of hazardous nature or resulting in decreased selectivity should be avoided. The common modifiers used in SFE are alcohols, which induce dipole/dipole interactions and hydrogen bonding with polar functional groups. Ethanol has been employed to increase the solubility of ginsenoids (Wang et al., 2001), antioxidants from tamarind seed coat (Luengthanaphol et al., 2004), Eucalyptus (El-Ghorab et al., 2003) and olive leaves (Tabera et al., 2004), lipophilic compounds from marigold (Baumann et al., 2004), and flavonoids and terpenoids from Ginko biloba (Yang et al., 2002). Methanol was used for the extraction of magnolol (Dean et al., 1998), soy isoflavones (Rostagno et al., 2002), phenolics from chamomille (Scalia et al., 1999), pistachio hulls (Goli et al., 2005) or aloe extracts (Hu et al., 2005). Isopropanol was used for the extraction of ginger oleoresin (Zancan et al., 2002). Water is the cosolvent employed in the industrial extraction of vanilla. The extracts resulting from the unselective extraction of compounds at high pressure can be fractionated in multiple-stage separators. The pressure and
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Table 18.10 Examples of extraction of plant food antioxidants using SFE Extraction yield (total extract, TE and active compounds)
Antioxidant activity1
Reference
Aloe vera leaf skin (Aloe barbadensis) Baical skullcap root (Scutellaria baicalensis) Boldo (Peumus boldus) Boldo (P. boldus) Coriander seeds (Coriandrum sativum) Eucalyptus leaf oil (Eucalyptus camaldulensis) Eucalyptus leaves (E. camaldulensis)
TF = 1.50 TF = 0.27
DPPH DPPH
TE = 2.9; B (0.0031)e TE = 4.9; B (7.4)e TE = 1.92 MTp H (0.242); STp (0.0825) pCl (1.16); oHC (0.17); MEP (0.34); Thy (0.98)a TE = 16.6
TEAC ABTS DPPH Inhibition of LA oxidation Inhibition of LA oxidation
Hu et al., 2005 Bimbato and Tamanini, 2003 Del Valle et al., 2005 Del Valle et al., 2005 Yépez et al., 2002 Fadel et al., 1999
Ginger (Zingiber officinale)
Monoterpenes, sesquiterpenes and gingerols
b-C, LA
Leal et al., 2003
Ginger (Z. officinale)
TE = 2.31
b-C, LA
Zancan et al., 2002
Helichrysum dried flower heads (Helichrysum italicum) Hop (Humulus lupulus) Lemon balm (Melissa officinalis subs. officinalis) (M. officinalis subs. inodora) Marjoram (Origanum vulgare) Hungarian marjoram (O. majorana) Egyptian marjoram Nigella seeds (Nigella sativa)
Flavonoids; TE = 4.5–4.9
O2–;
Marongiu et al., 2003
TP = 0.52–3.79/TF = 0.2–0.92 TE = 1.9 TE = 0.7
LA oxidation LA oxidation
Lermusieau et al., 2001 Marongiu et al., 2004
TE = 3.2 TE = 3.76 TE = 5.39 T.O. (1.14)b; Tq (0.49); p-Cy (0.23)
Peroxide value Inhibition of oil oxidation
Uy et al., 1991 Vági et al., 2005
b-C, LA
Machmudah et al., 2005
DPPH; b-C, LA
El-Ghorab et al., 2003
Extraction of natural antioxidants from plant foods 549
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Vegetal material (Latin name)
© Woodhead Publishing Limited, 2010
Vegetal material (Latin name)
Extraction yield (total extract, TE and active compounds)
Antioxidant activity1
Reference
Black pepper oleoresin (Piper nigrum)
RER = 5.43 (Pp: 39.4)
Tipsrisukond et al., 1998
Rosemary (Rosmarinus officinalis) Rosemary (R. officinalis)
TE = 5.2 TE = 7.15
Hexanal formation; TBARS Peroxide value
Rosemary (R. officinalis) Rosemary (R. officinalis)
Carnosic acid solubility TE = 4.5; Cm (39.6); Vbn (20.3); t-C (10.3)a
Rosemary (R. officinalis)
Sage (Salvia officinalis) Sage (S. officinalis)
F1: Rsl (3.6); Gk (11.9); Cr (5.6); CrA (66.0); MCr(1.6)a F2: Rsl (8.0); Gk (0.13); Cr (6.4); CrA (60.7); MCr(4.6)a TE = 1.6; Cm (0.025 ); Cn (0.001); CrA (0.236); RA(0.065)c TE = 5.7 TE = 5.02
Sage (S. officinalis)
TE = 46.26; F1 (024)/F2 (27.50)/F3 (18.52)
Rosemary (R. officinalis)
Black sesame seed (Sesamum indicum) Black sesame seed (S. indicum) Summer savory (Satureja hortensis)
TE = 51.83 Lignans (sesamin, sesamolin) S1: Crv (32); Lnl (24); Myc (22); Myl (21)c S2: Crv (44); Lnl (12); Myl (11); g-Tpnn (20)c S3: Crv (55); g-Tpnn (38)c Tamarind seed coat (Tamarindus indica) 0.13 mg EC g–1
b-C, LA Peroxide value; TBARS
Uy et al., 1991 Dapkevicius et al., 1998 Ramírez et al., 2005 Leal et al., 2003
b-C, LA DPPH
Cavero et al., 2005
b-C, LA
Carvalho et al., 2005
Peroxide value
Uy et al., 1991 Dapkevicius et al., 1998
AAb-c PF
Daukšas et al., 2001
DPPH; Lipid oxidation DPPH; lipid oxidation Inhibition oil oxidation
Hu et al., 2004 Xu et al., 2005a Esquível et al., 1999
Peroxide value
Luengthanaphol et al., 2004
550 Separation, extraction and concentration processes
Table 18.10 Continued
Tamarind seed coat (T. indica)
Thyme leaves (T. vulgaris) Thyme (T. vulgaris)
TE = 4.92 TE = 0.7; Thy (0.17)d
Turmeric oil (Curcuma longa) Turmeric (C. longa)
TE = 5.5 (STp H)
Green tea (Camelia sinensis)
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1
Zg A (40.1); Eg A (18.2); arT (15.7); Zg (3.1)
Peroxide value
Tsuda et al., 1995
Lipid oxidation
Chang et al., 2000
IPDEBTA Peroxide value
Ko et al., 2002 Uy et al., 1991 Dapkevicius et al., 1998
b-C, LA Rancimat Rancimat AOP
a
b-C, LA
Simándi et al., 2001 Zeković et al., 2001 and 2003 Gopalan et al., 2000 Leal et al., 2003
Antioxidant activity reported in the referenced work or, if not indicated, other works referencing antioxidant activity of the plant extracts. Relative percentage of the normalized area detected by LC–MS or GC. b Total oil (essential oil). c Relative weight content in the extract; d mg (100 g)–1. e mg g–1. TE, total extraction yield (%); TF, total flavonoids extraction yield (%); TP, total phenolics extraction yield (%); RER, relative extraction rate. AAb-c = antioxidant activity during bleaching of b-carotene–linoleate solution, scale (0:Low–5:High). AOP = antioxidant potency with linoleic acid in an iron/ascorbate system. b-C, LA = absorbance of the b-carotene–linoleic acid system. Hexanal = percentage of reduction in the formation of hexanal respect to control. IP DEBTA = inhibition of linoleic acid peroxidation by the DEBTA (diethyl-2-thiobarbituric acid) method. PF = protection factor calculated as the ratio between the induction period of the sample with additive and the control sample during vegetable oil oxidation. TBARS = percentage of reduction in the formation of TBARS with respect to control. TEAC = Trolox equivalent antioxidant capacity. O 2– = superoxide radical scavenger capacity of superoxide dismutase. Simple phenolics, derivatives and flavonoids: B, boldine; Crv, carvacrol; DPA, 3,4-dihydroxyphenyl acetate; EgA, E-g-atlantone; EC, epicatechin; ECG, epicatechin gallate; EGCG, epigallocatechin gallate; GA, gallic acid; Gk, genkwanin; HDA, 2-hydroxy-3¢,4¢-dihydroxyacetophenone; MDB, methyl 3,4-dihydroxybenzoate; MEP, 2-methyl-6-ethylphenol; RA, rosmarinic acid; Thy, thymol. Terpenoids: t-C, trans-caryophylene; Cm, camphor; Cn, cineole; Cr, carnosol; CrA, carnosic acid; p-Cy, p-cymene; MCr, methyl carnosate; Lnl, linalool; MTp H, monoterpenes hydrocarbons; Myc, myrcene; Myl, myrtenol; Rsl, rosmanol; STp, sesquiterpenes; STp H, sesquiterpene hydrocarbons; ar-T, ar-turmerone; g-Tpnn, g-terpinene; Vbn, verbenone; ZgA, Z-g-atlantone; Zg, zingiberene. Other compounds: pCl, p-cymen-7-ol; oHC, O-hydroxycumine; Pp, piperine; Tq, thymoquinone. a
Extraction of natural antioxidants from plant foods 551
Terminalia catappa leaves, seeds Thyme (Thymus vulgaris) Thyme (T. vulgaris)
TE = 0.29; HDA (0.043); MDB (0.068); DPA (0.885); EC (0.164)d EGC (290); EGCG (510); ECG (105); EC (90); GA (7)e Tannin, flavonoid glycosides TE = 2.0 TE = 5.46
552 Separation, extraction and concentration processes temperature in the separators influence the fractionation of extracts. Series of two or three separation stages are used to recover the desirable components. Although the extraction yields are increased under harsh conditions, the selectivity is reduced, leading to extracts with poor antioxidant activity. These instances require optimization of the extraction stage (Gelmez et al., 2009) and/or separation steps. The yield can be similar to the use of a stepwise increase in the extraction pressure, but with a lower consumption of solvent. 18.5.5 Other novel extraction technologies The fundamentals and application of alternative techniques using solvents at low pressure for the extraction of bioactive compounds have been comprehensively reviewed (Meireles, 2009). Some of these techniques have been developed for analytical purposes, but they could be scaled up. Studies dealing with the antioxidant activity of extracts obtained with ultrasonic- and microwave-assisted technologies are summarized in Table 18.11. Ultrasound-assisted extraction (UAE) is a cheap, scalable technique allowing reduced extraction times, decreased extraction temperature, and increased extraction yields. Its effects are mainly attributed to cavitation forces upon the propagation of the acoustic waves. Ultrasounds result in physical, chemical, and mechanical effects promoting the release of soluble compounds from the plant body, enhancing mass transfer and facilitating the access of solvent to cell contents (Knorr et al., 2002; Ma et al., 2008). UAE was used to extract valuable compounds such as phenolic acids and flavanone glycosides (Ma et al., 2008). In UAE, both treatment time and (particularly) temperature have significant effects on the properties of extracts. However, UAE must be applied carefully to avoid degradation of susceptible solutes (Ma et al., 2008). Degradation of phenolics depends on their substitution pattern. High comparative stability has been reported for sinapic and vanillic acids, which have methoxyl-type substituents in their aromatic rings, whereas benzoic acids are more stable than cinnamic acids. In previously dried samples, ultrasound accelerated rehydration and swelling, without significant chemical degradation (Qadir et al., 2009). The ultrasonic power had a positive effect on the contents of phenolic acids from mandarin peels (Ma et al., 2008), tannins from myrobalan nuts (Sivakumar et al., 2008), and anthraquinones from Morinda citrifolia roots (Hemwimol et al., 2006). Differences in the effects of acoustic treatments can be ascribed to the utilization of different equipment, effects of the nature of the solid matrix (hardness, compactness, and solute distribution), and cavitation. Some of the general trends outlined above were confirmed in the extraction of active components from Hypericum perforatum L. (Smelcerovic et al., 2006). Other technologies enhancing mass transfer in plant tissues and permeability of cytoplasmatic membranes include microwave-assisted extraction (MAE), pulsed electric fields (PEF) and high hydrostatic pressure (HHP) (Corrales
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Table 18.11 Other extraction technologies applied to the recovery of antioxidants from food plants Operational conditions
Extraction yield
Antioxidant activity
Reference
Ultrasound Almond shells
LSR 1:10; 5% NaOH; RT; 10 min
TEP (%)10.1
DPPH
Bearberry leaves
LSR 1:25; W:E; 60 kHz; 25 ºC; 40 min
2.15 %
Cow cockle seeds (Saponaria vaccaria)
LSR 1:50; (W, 50% E, ACN); RT; 1 h
W: 0.78 g GAE g–1 dw 50%E: 1.58 g GAE g–1 dw ACN: 0.20 g GAE g–1 dw
Dornfelder skins (Vitis vinífera) Ginseng (Panax ginseng) Ginseng (P. ginseng) Indian mulberry (Morinda citrifolia) Indian mulberry roots (M. citrifolia) Mandarin peels (Citrus unshiu)
35 kHz; 70 °C; 1 h
360 mmol GAE g–1 dw
Inhibition of oil oxidation DPPH ABTS FRAP ABTS
Ebringerová et al., 2008 Gribova et al., 2008
Corrales et al., 2008
LSR 1:50; 70% E; 60 ºC; 40 min; 250 W; 42 kHz LSR 1:50; 70% E; 70 °C; 2 h; 42 kHz; 284 W RLS 100; E; 60 °C; 60 min; 15.7 W
TS: 3.89 %
DPPH
Chen et al., 2010
G: 13.7 mg g–1 TP: 7.1 mg g–1 62.23%
DPPH
Kim et al., 2007
DPPH
79.62%
DPPH –
Pistachio (Pistachia vera) Plum (Prunus domestica)
LSR 1:8; W; M; 45 min
TCE:1935.12 mg g–1 dw FG: 1374.35 mg g–1 dw W: 34.2 mg TAE g–1 dw M: 32.8 mg TAE g–1 dw TP:1.74–3.75 mg GAE g–1 TP:1.18–2.37 mg CE g–1
Hemwimol et al., 2006 Pongnaravane et al., 2006 Ma et al., 2008
Peroxide value
Goli et al., 2005
ABTS
Kim et al., 2003
LSR 1:10; E; 60 ºC; 2 h; 38.5 kHz; 270 W LSR 1:80; 80% M; 30 °C; 10 min; 8 W
LSR 1:10; 80% M; -; 20 min
Güçlü-Üstündağ and Mazza, 2009
Extraction of natural antioxidants from plant foods 553
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Raw material
© Woodhead Publishing Limited, 2010
Raw material
Operational conditions
Extraction yield
Antioxidant activity
Reference
Sage (Salvia officinalis) Sea-buckthorn (Hippophae rhamnoides) Sempervivum (Sempervivun marmoreum) Wheat bran (Triticum aestivum) Microwave
LSR 1:4; M; RT, 20 min
–
DPPH
LSR 1:10; E; 30 °C; 60 min
15.6 (0.39) mg GAE g–1
DPPH TEAC
Ollanketo et al., 2002 Sharma et al., 2008
LSR 1:10; dW; 25 °C; 60 min; 40 kHz
2.5 mg g–1
Hydroxyl DPPH
Stojičević et al., 2008
LSR 1:10; 5%NaOH; 60 °C; 60 min; 100 W
24.3% polysaccharides
DPPH
Hromádková et al., 2008
Sappan wood (Caesalpinia sappan)
540 W; 20 min
0.747 g
DPPH Nitric oxide
Shrishailappa et al., 2007
Cranberry (Vaccinium macrocarpon)
LSR 1:5.7; 10 min; 125 °C W E A LSR 1:50; 70% M; 50 W; 10 min
Ginseng (Panax ginseng) Ginseng (Panax ginseng) Indian mulberry (Morinda citrifolia)
LSR 1:50; 70% E; 250 W; 10 min 2450 MHz LSR 1:100; E; 720 W; 30 min; 60 °C
17.6% (9.42 mg Q g–1) 2% (1272 mg Q g–1) 3.25% (960 mg Q g–1) G: 9.8 mg TAE g–1 dw TP: 5.8 mg TAE g–1 dw TS: 3.30% 65.88%
Inhibition of lipid Raghavan and oxidation Richards, 2007 TBARS DPPH
Kim et al., 2007
DPPH
Chen et al., 2010
DPPH
Hemwimol et al., 2006
554 Separation, extraction and concentration processes
Table 18.11 Continued
141 W; 83 s; 9.8 mL; 4 cycles
5.94 mg A mg–1
DPPH
Zhao et al., 2009
LSR 1:10; E 95%; 30 min; 500 W; 2450 MHz, 80 °C
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ACN, acetonitrile; E, ethanol; LSR, liquid:solid ratio; M, methanol; RT, room temperature; W, water; dW, distilled water. A, astaxanthin; CE, catechin equivalent; FG, flavanones glycosides; G, ginsenosides; GAE, gallic acid equivalent; Q, quercetin; TAE, tannic acid equivalent; TCE, total content of extract; TP, total phenolics; TEP, total extractable phenolics; TS, total saponins.
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Green algae (Haematococcus pluvialis) Longan peel (Dimocarpus longan Lour.) Teja (Cinnamomum iners Reinw.) Rice bran (Oryza sativa)
556 Separation, extraction and concentration processes et al., 2008). PEF enhances mass transfer by improving softness and texture (Eshtiaghi and Knorr, 2002), and increasing the contents of some solutes, including total phenolics. However, the antioxidant activity varies significantly with the treatment used (Corrales et al., 2008). PEF inactivates degrading enzymes. HHP increases the extraction yields by deprotonation of charged groups, disruption of salt bridges and hydrophobic bonds in cell membranes and decreasing the dielectric constant of water (Corrales et al., 2008). Combinations of the different extraction methods are used. High-pressure extraction combined with sonication was used to extract compounds from Berberis koreana bark, improving the extraction of bioactive compounds compared with the high-pressure process (Qadir et al., 2009).
18.6 Integration of extraction processes and purification Several technologies with different degrees of sophistication (including solvents, resins and membranes) were used to obtain purified compounds or fractions with biological activity from crude extracts. Some examples of their individual and combined application to the concentration and purification of natural antioxidants from several sources are included here. Combinations of conventional and emerging extraction technologies were proposed to extract and/or to purify natural compounds with antioxidant and biological activity. The extraction of oil and/or lipophylic fractions with supercritical fluids at low pressures and the subsequent solvent extraction of the solid residue was proposed to recover essential oils from aniseed (Rodrigues et al., 2003), Melissa officinalis (Marongiu et al., 2004), rosemary, sage (Nakatsu and Yamasaki, 2000), summer savory (Esquível et al., 1999), and oils from borage seeds (Soto et al., 2008). A sequence of SFE and hydrothermal extraction has been proposed to obtain antioxidants from bamboo (Quitain et al., 2004). Alternatively, a first extraction of the raw material with conventional solvents and further purification of the crude extracts by SFE offer operational advantages, including reduction in CO2 consumption, increased extraction yields, and processing of less substrate owing to its enhanced content of bioactive compounds. Two options are used to purify the bioactive compounds from the crude extract: relatively mild SFE to extract non-polar compounds (desirable when antioxidant compounds of a phenolic nature have to be recovered from the residue) (Hadolin et al., 2004), and SFE under harsh conditions to recover polar compounds in the extract (Díaz-Reinoso et al., 2006). A simple purification process based on the sequential extraction of the raw materials or extracts with solvents of different polarity was applied to the extraction of antioxidant compounds from plant foods. Cheung et al. (2003) reported on the processing of edible mushrooms, leading to similar yields as those obtained with methanol and water on large and small scales. The
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Extraction of natural antioxidants from plant foods 557 chemical components of the methanol extract were fractionated by solvents of increasing polarity (dichloromethane, ethyl acetate, butanol and water) (Cheung and Cheung, 2005). The dichloromethane subfraction of the methanol extract of Volvariella volvacea had a high antioxidant activity against lipid peroxidation of rat brain homogenate, whereas the ethyl acetate subfraction of the methanol extract of V. volvacea has antioxidant activity against the oxidation of human low density lipoprotein (LDL) similar to that of caffeic acid. Combined extraction involving pressurized hot water extraction as a first stage has been reported for a variety of purposes (Chambers et al., 1984; Wang and Weller, 2006). The advantages of membrane processing include a low energy requirement, no additives, mild operating conditions, separation efficiency and easy scaling up. The utilization of membrane technologies for concentrating and purifying bioactive phenolic compounds from aqueous streams is a topic of growing interest. For example, fractionation of phenolics into low- and high- MW compounds was reported for conventional solvent crude extracts from edible mushrooms (Cheung and Cheung, 2005), grape (Nawaz et al., 2006), grape pomace (Díaz-Reinoso et al., 2009), partially purified extracts from persimmon pulp (Ga et al., 2008), and mulberry root cortices (Yu et al., 2007). Ultrafiltration (UF) membranes are easy and fast to use in the separation of phenolics according to their MW. At the analytical scale, almond skin phenolics were fractionated to separate low MW compounds in permeate and proanthocyanidin oligomers (up to decamers) in retentate. Permeates from a 50 kDa membrane contained proanthocyanidin pentamers (Prodanov et al., 2008). UF membranes can be used for tailoring grape anthocyanins (containing about 60% monomers, 20% polymers, and 20% other forms) from grapes, having benzoic acids, hydroxycinnamates, anthocyanins, flavan-3-ols, and flavonols as major phenolic groups. The membrane (<100 kDa) separated the polymeric form in the retentates, wheareas permeates contained low-MW compounds. After a 10-fold reduction in volume, the permeate had higher monomer and lower polymer content than the initial feed, with the opposite for the retentate. The composition of the fractions affected color, functional properties and antioxidant activity. The total phenolic content correlated linearly with the antioxidant activity, whereas lightness and color properties related linearly to the monomeric content (Kalbasi and Cisneros-Zevallos, 2007). Ultrafiltration (0.22 mm) of 50% ethanol solvent of grape seeds was employed to maximize the recovery of polyphenols (Nawaz et al., 2006), whereas nano- and ultrafiltration were used to process aqueous extracts from grape pomace (Díaz-Reinoso et al., 2009). Membrane separation was applied to the extracts from both conventional and alternative extraction processes. Samples of edible mushrooms (Lentilula edodes, V. volvacea) were extracted sequentially with petroleum ether, ethyl acetate and methanol, and the insoluble residue was first extracted with
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558 Separation, extraction and concentration processes boiling water and further processed by UF (10 kDa) to yield fractions of low and high MW. The low-MW fraction was highly active against lipid peroxidation of rat brain homogenate (Cheung and Cheung, 2005), and the activity seemed to correlate with their content of protein/free amino acids. SFE, ultrasound-aided extraction and membranes were applied to the manufacture of algal polysaccharides. Sargassum pallidum powder was extracted by SFE and the resulting powder was disrupted by ultrasonic waves, incubated, and centrifuged. Proteins in the supernatant solution were removed by adding trichloroacetic acid, and the resulting solution was adjusted to pH 7, passed through membranes having different MW cut-offs, and subjected to rotary evaporation, ethanol precipitation and freeze-drying. Crude polysaccharides obtained from UF were dissolved and applied to a column of DEAE-52 cellulose to obtain seven fractions of polysaccharides (Ye et al., 2008). Among them, the low-MW fractions with high sulfate content presented the highest antitumor activity against HepG2 cells, A549 cells, and MGC-803 cells (Ye et al., 2008). SFE was used for purification of tea polyphenols as part of a complex sequence involving ultra- and nanofiltration (Fig. 18.3), vacuum concentration, spray drying and conventional solvent extraction. A 1000-Da membrane was used to process spinach extract produced by pressurized liquid water or by ethanol extraction, yielding a high-MW fraction with much higher antioxidant activity than that of the low-MW fraction (Howard and Pandjaitan, 2008). The separation selectivity is improved by using membranes able to interact with some compounds in solution. For example, a modified PVP membrane favoring the formation of hydrogen bonds with flavonoids improved their separation from G. biloba extracts (Xu et al., 2005b). UF was used in concentrating melanoidins from coffee brew. Operating in discontinuous diafiltration mode (Rufián-Henares and Morales, 2007), the retentates contained soluble melanoidins with antioxidant activity, whereas permeates contained compounds non-covalently bound to melanoidins, which showed no antioxidant activity (Delgado-Andrade and Morales, 2005). Novel technologies (such as membrane distillation and osmotic evaporation) were proposed to produce juice concentrates preserving their natural quality and antioxidant activity (Cassano et al., 2007; Koroknai et al., 2008) (Fig. 18.4). The combination of protease hydrolysis and membrane fractionation of peptides is an attractive method to improve the potential of proteins with high nutritive value but poor functional properties (e.g. solubility, emulsification and film-forming properties) and negative sensory properties (color, taste and texture). The functional and immunological properties of proteins can be modified, enabling their utilization in applications such as additives for beverages, hypoallergenic infant diets, nutritional therapies, food texture enhancers, or pharmaceutical ingredients. The functionality of the final product can be controlled by selecting specific enzymes and reaction factors (Cheryan
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Fig. 18.3 Processing scheme involving supercritical CO2, and ultra- and nanofiltration.
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Extraction of natural antioxidants from plant foods 561 and Deeslie, 1983; Deeslie and Cheryan, 1988). Fractionation of protein hydrolyzates is a typical area for membrane applications. Examples dealing with conventional and non-conventional protein sources are presented. Rapeseed protein isolates are suitable for food applications owing to their functional properties and health benefits, such as ACE inhibition, bile acidbinding and free radical-scavenging activity. Since the antinutritive and toxic compounds (such as glucosinolates and phytates) have significantly lower MW than rapeseed proteins, they can be selectively separated by precipitation at controlled pH or removed by UF. The ultrafiltered protein extract had good functional properties, whereas the precipitated protein showed stronger ACE inhibition, bile acid-binding capacity and DPPH radical-scavenging capacity (Yoshie-Stark et al., 2008). Soybean peptides, with or without enzymatic treatment, were selectively retained by UF membranes (Moure et al., 2006) (Fig. 18.5). Alfalfa leaf protein was hydrolyzed with protease, and the reaction products, fractionated by UF and purified by adsorption, showed high nutritive value, chelating ability, reducing power, and radical scavenging activity. Mouse breeding with these products resulted in increased activities of glutathione peroxidase and superoxide dismutase, and decreased the formation of malonaldehyde by oxidative reaction (Xie et al., 2008). Proteins extracted from potato tubers and by-products from the potato industry were hydrolyzed and ultrafiltered to yield a permeate containing ACE-inhibitory compounds (Pihlanto et al., 2008). Papain was used for preparing hydrolyzates from wheat gluten, which were separated by UF to yield fractions with strong antioxidative activities measured by the linoleic acid and DPPH tests (Wang et al., 2007). The hydrolyzates produced by enzymatic treatment of wheat gluten with commercial proteases were fractionated by UF, to yield products with properties strongly related to their molecular weight distribution and amino acid composition. The activities of selected fractions were measured by the linoleic acid and scavenging radical assays (Kong et al., 2008). UF membrane reactors have several advantages over conventional technologies, including enzyme reuse and accurate selection of the product molecular size. The flowsheet of a process for producing soy protein hydrolyzate using a continuous UF membrane reactor fed with a combination of commercial enzymes is shown in Fig. 18.6. The amino acid sequences of peptides depended on the specificity of enzymes, whereas the characteristics of the membrane governed the functional properties of protein hydrolyzates (Chiang et al., 1999). The major polysaccharides from olive fruits are pectins, which are esterlinked to phenolic acids. The action of selected enzymes facilitates the release of phenols. Bouzid et al. (2005) proposed the incubation with coumaroyl ester hydrolase and further processing to recover hydroxytyrosol and free phenolic acids (caffeic, p-coumaric and ferulic acids). Adsorption onto polymeric resins was successfully used for the separation
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and purification of phenolic compounds from natural products (Kammerer et al., 2005; Llorach et al., 2004; Saleh et al., 2008; Scordino et al., 2005). Separation and purification of rutin from leaves, flowers, stalks, and seeds of buckwheat (Fagopyrum esculentum Moench) has been reached by a lowcost process consisting of extraction with aqueous 50% ethanol at 80 °C for 1 h, adsorption onto a styrene-based resin (HP-20), elution with water and aqueous ethanol (20% and 30%), and recrystallization at 4 °C for 12 h). Rutin recovery was 92%, with over 95% purity (Kyoung et al., 2005). Extraction of dried apple pomace (a by-product of apple juice production) with diluted mineral acid, followed by adsorption of phenolics (phloridzin, chlorogenic acid and quercetin glycosides) was carried out using a hydrophobic styrene–divinylbenzene resin (Amberlite XAD 16HP). Pectins were eluted with distilled water, and the phenolic compounds with methanol (Schieber et al., 2003). Water and alcoholic extracts from lettuce, chicory by-products and cauliflower (Brassica oleracea L. var. botrytis) were purified by adsorption on Amberlite XAD-2. The antioxidant capacity (reducing power and DPPH, ABTS radical scavenging capacity) and the capacity to inhibit lipid peroxidation (ferric thiocyanate assay) were linearly correlated with the phenolic content (Llorach et al., 2003 and 2004). Polysaccharide fractions prepared from various materials are usually prepared by sequential steps of solubilization, precipitation and fractionation or purification. For example, crude extracts from Cuscuta chinensis seeds were extracted with hot water and diluted alkali, and then precipitated by addition of ethanol. Anion exchange and gel filtration chromatography were used for purification of an acidic polysaccharide with rhamnogalacturonanlike structure (Bao et al., 2002). Alkaline extraction of ferulic acid from © Woodhead Publishing Limited, 2010
564 Separation, extraction and concentration processes maize bran was carried out using 2M NaOH. Ferulic acid was purified by adsorption chromatography followed by preparative high-performance thinlayer chromatography, yielding a product with a purity of 95.35% (Tilay et al., 2008). Novel adsorbents for polyphenols include silk fibroin, which has been used to process ethanol extracts from olive leaves. This material adsorbed 15 mg g–1 rutin and 96 mg g–1 oleuropein, as well as other polyphenols such as verbascoside, apigenin-7-glucoside, and luteolin-7-glucoside (Altiok et al., 2008). Woody materials such as agricultural by-products or natural fibres are useful for isolating bioactive components from plants. Substrates prepared from woody tea stalk, pine sawdust and sugarcane bagasse were suitable to adsorb decaffeinated catechins from tea extracts. The tea components were separated by gradient elution with increasing ethanol concentration, yielding fractions with different proportions of caffeine, partial non-gallated catechins and gallated catechins. The concentration of total catechins in the selected ethanol eluates was above 90%, with caffeine below 1% (dry basis) (Ye et al., 2009). The water-soluble fraction resulting from steam treatment of olives (Fernández-Bolaños et al., 2002) contains other high-added-value compounds such as monosaccharides, oligosaccharides, and mannitol. Oligosaccharides were separated by size-exclusion chromatography, whereas highly purified mannitol could be recovered by a simple purification method (FernándezBolaños et al., 2004). Highly purified 3,4-dihydroxyphenyl glycol (DHPG) was obtained from alperujo by chromatography to yield 96% pure product at 21% yield. From 1000 kg of wet alperujo (300 kg of dry matter), 807 g DHPG could be obtained (Rodríguez et al., 2009). Hydroxytyrosol was recovered from olive cake at good yield (1–1.2 g hydroxytyrosol/100 g of dry matter) upon acidic processing (Fernández-Bolaños et al., 2002). Crude rapeseed peptides and peptide fractions with antioxidant and antithrombotic activities were prepared by incubation with enzymes and further adsorption. The rapeseed slurry from a wet-milling was treated with a combination of enzymes, and subjected to thermal inactivation and centrifugation. The aqueous phase was adjusted to pH 4 with acetic acid, adsorbed in the macroporous adsorption resin and eluted with deionized water and ethanol. Stepwise desorption with 25–55% ethanol enabled the recovery of different fractions. The fraction eluted with 55% ethanol showed more potent antioxidant activities (reducing power, inhibition of lipid oxidation in a liposome model) except for hydroxyl radicals, probably owing to the higher contents of hydrophobic amino acid, tannin, and the brown color substances. A positive correlation existed between the peptide concentration and antioxidant activity (Zhang et al., 2007). Tricin (5,7,4¢-trihydroxy-3¢,5¢-dimethoxyflavone), a cancer chemopreventive agent, was prepared from a concentrated extract from bamboo leaves and further processed by adsorption in a polystyrene resin, preparative high-
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Extraction of natural antioxidants from plant foods 565 performance liquid chromatography, dialysis and crystallization (Jiao et al., 2007), resulting in higher yields and fewer processing steps than the normal synthetic approach. Membranes and resins were proposed for processing hydrolyzates from alkaline treatments. Liquors from the alkaline processing of sugarcane bagasse were ultrafiltered, and the phenolic acids present in permeates were bound to an anion-exchange resin. After washing, elution with a mixture of water–ethanol–HCl followed by crystalization and washing of crystals with 1% HCl led to a concentrate containing products with 89.7% purity (measured as coumaric acid), which had the same antioxidant activity, reducing power and free radical scavenging capacity as the standard p-coumaric acid (Ou et al., 2009). Combinations of different treatments were also applied to the extraction and purification of polysaccharides from Physalis alkekengi var. francheti fruit, including hot water extraction, ultrasonic-assisted extraction and enzyme pretreatment. Fractionation of polysaccharides with DEAE and Sephadex G-200 led to fractions with different radical scavenging activity (Ge et al., 2009). Membranes and resins have been combined in processing schemes similar to that given in Fig. 18.7. A macroporous adsorption resin was used to remove low-MW polar substances (including sugars, gallic acid, and organic acids) and to increase the contents of total polyphenols and condensed tannins. The partially purified extract was subjected to UF (10 kDa) to yield a retentate containing condensed tannins of high MW, whereas permeate contained lowMW phenolic compounds. High-MW tannins showed high hydroxyl radical scavenging activity, as a result of their structure. This fraction also exhibited dose-dependent, inhibitory activity against hydroxyl and superoxide anion radical scavenging activity, as well as activity in the peroxidation of linoleic acid. Low-MW tannins were reported to be pro-oxidant during accelerated lipid peroxidation (Gu et al., 2008). UF of green tea water extracts through a composite membrane was suitable for retaining particles, proteins, polysaccharides and tannic acid, leading to a permeate containing more than 40% phenols (epicatechin, epicatechin gallate, epigallocatechin, epigallocatechin gallate) and caffeine (Li et al., 2005). This stream was further separated using adsorption–desorption resins, to yield a purified product containing more than 90% polyphenols (Li et al., 2005). A combination of membrane fractionation and resin purification was also applied to peptide purification. A recent example is the purification of nucleoprotein complexes isolated from Saccharomyces cerevisiae by mild alkaline extraction and precipitation with acetic acid. The high- and low-MW fractions of the nucleoprotein were separated by cross-flow microfiltration (Butylina et al., 2007). S. cerevisiae cells were exposed to hydrogen peroxide or to ultraviolet irradiation to induce oxidative∑ stress: the first acts as an oxidant by producing hydroxyl free radicals (OH ) that attack cell components and cause various modifications in human cells, whereas the latter can cause
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Extraction of natural antioxidants from plant foods 567 damage to nucleic acids and proteins, and affect other molecules through production of ROS. The addition of a high-MW nucleoprotein complex to damaged cells enhanced growth rate of the yeast population up to a similar value to that of intact cells (Butylina et al., 2007).
18.7 Future trends Antioxidants are commonly used in foods to retard or prevent deterioration via lipid oxidation, which leads to the development of undesirable rancid odors, off-flavors, discoloration and generation of potentially toxic compounds, limiting the quality, acceptability and shelf life of processed products. As a general trend, obtaining a similar degree of protection as synthetic antioxidants requires higher doses of natural extracts. However, if the extracts lack toxicity and add functional and biological properties to the product, higher loadings could be used. The use of natural antioxidants would be better restricted to those cases in which their application is necessary, and the use of materials relatively rich in antioxidants should be preferred to the application of extracts, concentrates or pure mixtures of active components. In this context, the utilization of food processing wastes as feedstocks could facilitate the production of valuable natural products, which would guarantee both sustainability and satisfaction of consumer demands. In the future, the application of advanced, efficient technologies such as ultrasound, pressurized extraction, and SFE is expected to offer great potential and selectivity for process development. Technological advances in these methods would require specific optimization. Compared with conventional extraction methods, the combination of low cost raw materials and effective extraction technologies is of environmental and economical interest. These techniques do not involve the utilization of organic solvents, allow reduced extraction times and may improve separation selectivity.
18.8 Sources of further information and advice Owing to the wide-ranging nature of this work, the set of references included provide a non-exhaustive overview of the field. Because of this, special attention has been devoted to include recent reviews, dealing with specific topics, which can be of help for futher information. However, the intensive research on antioxidants is reflected in a high rate of publication, particularly in topics related to the bioactivity of selected compounds suitable for key applications such as dermatological preparations, cancer, cardiovascular and neurodegenerative diseases.
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18.9 Acknowledgements The authors are grateful to the Spanish Ministry of Education and Science (MEC) for the financial support of this work (in the framework of the Research Project reference ALG2006-05387, which had partial financial support from the FEDER funds of the European Union).
18.10 References Adil I H, Yener M E and Bayindirli A (2008), ‘Extraction of total phenolics of sour cherry pomace by high pressure solvent and subcritical fluid and determination of the antioxidant activities of the extracts’, Separation Science and Technology, 43, 1091–1110. Aehle E, Raynaud-Le Grandic S, Ralainirina R, Baltora-Rosset S, Mesnard F, Prouillet C, Mazière J-C and Fliniaux M-A (2004), ‘Development and evaluation of an enriched natural antioxidant preparation obtained from aqueous spinach (Spinacia oleracea) extracts by an adsorption procedure’, Food Chemistry, 86, 579–585. Ahn G N, Park E J, Kim D S, Jeon Y J, Shin T K, Park J W, Woo H C, Lee K W and Jee Y (2008), ‘Anti-inflammatory effects of enzymatic extract from Ecklonia cava on TPA-induced ear skin edema’, Food Science and Biotechnology, 17, 745–750. Ahn M-J, Yoon K-D, Min S-Y, Lee J S, Kim J H, Kim T G, Kim S H, Kim N-G, Huh H and Kim J (2004), ‘Inhibition of HIV-1 Reverse transcriptase and protease by phlorotannins from the brown alga Ecklonia cava’, Biological & Pharmaceutical Bulletin, 27, 544–547. Almeida I F, Valentão P, Andrade P B, Seabra R M, Pereira T M, Amaral M H, Costa P C and Bahia M F (2008), ‘In vivo skin irritation potential of a Castanea sativa (chestnut) leaf extract, a putative natural antioxidant for topical application’, Basic & Clinical Pharmacology & Toxicology, 103, 461–467. Alonso-Salces R M, Korta E, Barranco A, Berrueta L A, Gallo B and Vicente F (2001), ‘Determination of polyphenolic profiles of Basque cider apple varieties using accelerated solvent extraction’, Journal of Agricultural and Food Chemistry, 49, 3761–3767. Altiok E, Baycin D, Bayraktar O and Ülkü S (2008), ‘Isolation of polyphenols from the extracts of olive leaves (Olea europaea L.) by adsorption on silk fibroin’, Separation and Purification Technology, 62, 342–348. Altunkaya A, Becker E M, Gökmen V and Skibsted L H (2008), ‘Antioxidant activity of lettuce extract (Lactuca sativa) and synergism with added phenolic antioxidants’, Food Chemistry, 115, 163–168. Amaral S, Mira L, Nogueira J M F, da Silva A P and Florencio M H (2009), ‘Plant extracts with anti-inflammatory properties – a new approach for characterization of their bioactive compounds and establishment of structure – antioxidant activity relationships’, Bioorganic & Medicinal Chemistry, 17, 1876–1883. Amarowicz R and Shahidi F (1997), ‘Antioxidant activity of peptide fraction of capelin protein hydrolysates’, Food Chemistry, 58, 355–359. Ames B N, Shigenaga M K and Hagen T M (1993), ‘Oxidants, antioxidants, and the degenerative diseases of aging’, Proceedings of the National Academy of Sciences of the United States of America, 90, 7915–7922. Andrade P B, Pereira D M, Ferreres F and Valentao P (2008), ‘Recent trends in high throughput analysis and antioxidant potential screening for phenolics’, Current Pharmaceutical Analysis, 4, 137–150. Andreasen M F, Christensen L P, Meyer A S and Hansen A (1999), ‘Release of hydroxycinnamic and hydroxybenzoic acids in rye by commercial plant cell wall
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Fractionation of egg proteins and peptides 595
19 Fractionation of egg proteins and peptides for nutraceutical applications B. P. Chay Pak Ting, Y. Pouliot and S. F. Gauthier, Laval University, Canada and Y. Mine, University of Guelph, Canada
Abstract: The main components of egg proteins and their physicochemical characteristics, and bioactive peptides derived by enzymatic hydrolysis from egg proteins, are described. An overview of recent developments in fractionation and purification processes for bioactive proteins/peptides from egg proteins is also presented. Further developments of techniques required to achieve the separation of specific proteins/peptides are described and the potential for creating new value-added ingredients with applications in the food, nutraceutical and biotechnological industry is explored. Key words: egg proteins, proteins, peptides, separation, bioactivity.
19.1 Introduction The avian egg is considered to be a rich source of nutrients, such as proteins, lipids, and enzymes, and biological substances, such as growth promoting factors and defence factors. Overall, an egg is constituted of 63% egg white, 27.5% egg yolk and 9.5% eggshell (Table 19.1). Although eggs contain about 75% water, they are a rich source of high-quality protein, unsaturated fatty acids, vitamins and minerals. Egg proteins are essentially distributed between the egg white and the yolk, with a small proportion in the eggshell. Lipids, negligible in egg white, are almost exclusively found in the egg yolk and are associated with proteins to form lipoproteins. Carbohydrates are minor egg components, which are present in the egg as both free carbohydrates and bound to proteins/lipids. Most of the minerals are found in the eggshell and in the yolk, where phosphorus and potassium are in soluble form.
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596 Separation, extraction and concentration processes Table 19.1 Chemical composition of hen eggs and their major components [data from Li-Chan and Nakai (1989)] Constituent
Egg shell Egg white Egg yolk
% (w/v)
9.5 63.0 27.5
Major components (%, w/w) Protein
Lipid
Carbohydrate
Ash
6.4 9.7–10.6 15.7–16.6
0.03 0.03 31.8–35.5
– 0.4–0.9 0.2–1.0
– 0.5–0.6 1.1
Table 19.2 Composition of egg white and yolk proteins
Egg white
% of total proteins
Ovalbumin Ovotransferrin Ovomucoid Ovomucin Lysozyme G2 globulin G3 globulin Ovoinhibitor Ovoglycoprotein Ovoflavoprotein Ovomacroglobulin Cystatin Avidin
54 12 11 3.5 3.4 4.0 4.0 1.5 1.0 0.8 0.5 0.05 0.05
Lipovitellin Phosvitin LDLg Livetin LDL
70 16 12 15 85
Egg yolk Granules Plasma
LDLg, low-density lipoprotein from granules.
Table 19.2 lists the major proteins of egg white and egg yolk together with their average concentration. The major proteins in egg white are ovalbumin, ovotransferrin, ovomucoid, ovomucin and lysozyme which account for >83% of total egg white proteins. The major egg yolk proteins take the form of lipoprotein complexes, which are divided into the plasma and granule fractions comprising lipovitellin, phosvitin, livetin and lipovitellenin. A number of egg proteins and some of their peptide sequences that have biological activities have been identified. For example, lysozyme is the egg protein that has probably attracted the most attention because of its potential use as an antimicrobial agent in foods. Chicken lysozyme contains a peptide sequence that is potently antimicrobial against both Gram-positive and Gram-negative bacteria. Biologically active proteins and peptides occur at relatively low concentrations in foods and pre-concentration of purification © Woodhead Publishing Limited, 2010
Fractionation of egg proteins and peptides 597 is needed to obtain a dose level to produce beneficial effects in situ. It was therefore necessary to develop technologies for fractionation and purification of bioactive molecules of interest for nutraceutical applications. This chapter mainly focuses on the chemistry and biological activities of proteins and peptides derived from egg components and the various technologies applied for fractionation/purification of proteins/peptides.
19.2 Composition and physicochemical characteristics of egg proteins and applications in the nutraceutical industry 19.2.1 Egg white The physicochemical characteristics of some egg white proteins are listed in Table 19.3. Egg white proteins are predominantly globular proteins having an acidic isoelectric point (pI), the exceptions being lysozyme and avidin. Except for lysozyme, egg white proteins are glycoproteins that are rich in sulfur amino acids and very sensitive to heat denaturation. Ovalbumin Ovalbumin has a molecular mass of 45 kDa and is composed of 385 amino acids. This protein belongs to the serpins family and possesses two genetic variants which differ at residue 289 and 311 by a substitution of Glu and Asn, respectively, by Gln and Asp (Ishihara et al., 1981; Wiseman et al., 1972). The primary sequence reveals one disulfide bond between Cys74 and Cys121. Table 19.3 Physicochemical characteristics of egg white proteins [adapted from LiChan and Nakai (1989)]
Amino acid residues
MW (kDa)
pI
T d*
Number of cystein residues [disulfide bridge]
Ovalbumin Ovotransferrin Ovomucoid Ovomucin Lysozyme G2 globulin G3 globulin Ovoinhibitor Ovoglycoprotein Ovoflavoprotein Ovomacroglobulin Avidin Cystatin
385 686 186 8300 129 – – – – 219 – 128 –
45 77.7 28 5500–8300 14.3 30–45 – 49 24.4 32 760–900 68.3 12.7
4.5 6.0 4.1 4.5–5.0 10.7 5.5 4.8 5.1 3.9 4.0 4.5 10 5.1
84 61 79 – 75 92.5 – – – – – 85 –
6 [1] 30 [15] 18 [9] – 8 [4] – – – – 18 [9] – 2 [1] 4 [2]
(–) Data not available. * Denaturation temperature.
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598 Separation, extraction and concentration processes Approximately one half of ovalbumin amino acids residues are hydrophobic and one third are acidic, which confer a pI of 4.5 (Li-Chan and Nakai, 1989). The secondary structure of ovalbumin consists of 30% a-helix and 32% b-sheet structure using high resolution x-ray diffraction (Stein et al., 1990, 1991). Three components A1, A2 and A3 have two, one and no phosphate, respectively, and are found in purified ovalbumin (Perlman, 1952). In solution, ovalbumin can be denatured and aggregated by thermal denaturation or by exposure to the air–water surface (Mine et al., 1990). Partial thermal denaturation of ovalbumin occurs at 78 to 86 °C. The protein adopts a more stable form, termed S-ovalbumin to denote its increased stability. At a heating rate of 10 °C min–1 at pH 9, the denaturation temperature of ovalbumin is 84.5 °C compared with 92.5 °C for S-ovalbumin (Donovan and Mapes, 1976). The greater stability, compactness and hydrophobicity of the S-form contrast with that of ovalbumin (Nakamura and Ishimaru, 1981). Ovotransferrin Ovotransferrin, also called conalbumin, belongs to the transferrin family. Ovotransferrin, with a molecular mass of 77.7 kDa and 686 amino acids, contains no phosphorus or free sulfhydryl group. Ovotransferrin has a pI of 6.0 and has the capacity to bind two Fe3+ ions per molecule with two CO2–3 or HCO3– ions (Mine, 2007). Ovotransferrin is folded into two lobes which show similar structural elements (Williams et al., 1982). Each lobe contains a site for iron binding and is further divided into two distinct domains. Ovotransferrin is the most heat-sensitive protein of egg white. The protein easily binds metallic ions, such as Fe3+, Al3+, Cu2+ or Zn2+ forming heatstable complexes. Ovomucoid Ovomucoid is a glycoprotein with a molecular weight of 28 kDa and a pI of 4.1. Its amino acid sequence is composed of 186 residues and it possesses no tryptophan residues. Ovomucoid has nine disulfide bridges and no sulfhydryl groups. Ovomucoid has three domains defined by the amino acid sequences of 1–68, 69–130 and 131–186, each domain is cross linked by three disulfide bridges (Kato et al., 1987). Its secondary structure is composed of 26% a-helix, 46% b-structure, 10% b-turn and 18% random coil structure (Watanabe et al., 1981). Ovomucoid is highly resistant to heat owing to its high cystine content and, consequently, to the high number of disulfide linkages. Under acidic conditions ovomucoid can resist to heat treatments up to 100 °C, but it is rapidly (89 °C) denaturated in alkaline solutions (pH 9) (Matsuda et al., 1982). Ovomucin Ovomucin is a highly glycosylated protein displaying a molecular weight in the range of 5500–8300 kDa depending on the method of isolation and environmental conditions (Li-Chan and Nakai, 1989). Ovomucin differs from
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Fractionation of egg proteins and peptides 599 other egg white proteins because it is extremely large. Its structure comprises sulfate esters, large amounts of cystine and 50% of the total sialic acid contents of egg white (Stadelman and Coterril, 1973). Ovomucin is composed of two subunits: a- and b-ovomucin, which differ by a carbohydrate-poor fraction (15% of dry matter) and a carbohydrate-rich fraction (50% of dry matter), respectively, and amino acids composition. Watanabe et al. (2004) estimated the molecular weight of a-ovomucin at 254 kDa whereas that of b-ovomucin has been estimated, using SDS-PAGE, to be between 400 and 720 kDa (Itoh et al., 1987). Ovomucin is heat stable over the pH range between 7.1 and 9.4. Cunningham and Lineweaver (1965) showed that the viscosity and absorbance of ovomucin solutions did not change upon heat treatment at 90 °C for 2 h. Lysozyme Lysozyme is a small enzyme molecule of 14.3 kDa capable of hydrolyzing specific polysaccharides that cleaves the b 1-4 linkages between N-acetylneuraminic and N-acetylglucosamine in bacteria cell. The single polypeptide chain consisting of 129 amino acid residues is cross-linked by four disulfide bridges. Because of its strong basic character (pI = 10.7), lysozyme binds to ovomucin, ovotransferrin and ovalbumin, via electrostatic interactions between positively charged lysozyme and negatively charged residues of sialic acid in these glycoproteins. Young et al. (1994) analyzed the three-dimensional structure of hen egg white lysozyme. The lysozyme molecule has two domains: the N-terminal domain (residues 40 to 88) with a hydrophobic core is composed of antiparallel b-sheets whereas the second hydrophilic domain, residues 1 to 39 and 89 to 129, is made up of a-helix. The enzyme is much more heat sensitive in egg albumen than when present alone. Minor egg white proteins The remaining proteins listed in Table 19.2 and 19.3 (G2 and G3 globulin, ovoinhibitor, ovoglycoprotein, ovoflavoprotein, ovomacroglobulin, cystatin, avidin) account for <17% of egg white proteins. Although some of these proteins may have potential applications as bioactive proteins or peptides, their sequence and structure and properties have not been fully characterized to date. 19.2.2 Egg yolk Egg yolk is a complex mixture of micro granules held in suspension. Proteins and lipids are the main constituents of yolk accounting for 15.7–16.6 and 31.8–35.5%, respectively. The yolk lipid fraction contains 66% triacylglycerol, 28% phospholipid, 5% cholesterol and minor amounts of other lipids (Powrie and Nakai, 1985). Egg yolk can be separated into plasma (the supernatant) and granule (the precipitate) after dilution in a saline solution (0.17 M
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600 Separation, extraction and concentration processes NaCl) followed by a centrifugation. Plasma is composed of 85% low-density lipoproteins (LDL) and 15% livetin and granules contain mainly 70% lipovitellin (HDLs), 16% phosvitin and 12% LDLg (Burley and Cook, 1961). HDLs and phosvitin come from the same precursor, namely vitellogenin II, and they are linked by phosphocalcic bridges which probably form the basic unit of the granules (Radomski and Cook, 1964). After uptake by the oocyte, vitellogenin undergoes specific enzymatic cleavage to generate lipovitellin I (120 kDa) and lipovitellin II (30 kDa) in the N-terminal and C-terminal region, respectively, with the phosphoseryl-rich domain termed phosvitin lying in between. Protein composition of egg yolk granules Phosvitin Egg yolk phosvitin is a phosphoglycoprotein with a molecular weight of 35 kDa containing 10% phosphorus and 6.5% carbohydrates (Mecham and Olcott, 1949). Phosvitin is constituted of 217 amino acid residues which comprise a core region of 99 amino acids, consisting of 80 serines, grouped in runs of maximally 14 residues interspersed by arginines, lysines and asparagines. Most of serine residues are phosphorylated and the phosphoserines are forming blocks that can carry up to 15 consecutive residues (Byrne et al., 1984; Van Het Schip et al., 1987). The relative abundance of phosphoseryl groups in the phosvitin amino acid sequence confers to the protein a large central hydrophilic portion surrounded by two small hydrophobic parts at the N-terminal and C-terminal. Owing to its polyanionic character (pI = 4), phosvitin possesses a very strong metal-chelating property. It can bind multivalent metals and 95% of Fe in egg yolk is complexed together with phosvitin (Greengard et al., 1964). Castellani et al. (2004) have found that pH 6.5 and ionic strength of 0.15 M were optimal for iron binding by phosvitin. Fourier transform infrared spectroscopy showed that the secondary structure of phosvitin is composed of 0% a-helix, 50% b-sheets, 7% b-turns and 43% random coil (Losso et al., 1993). Two constituents from hen’s egg yolk phosvitin, namely a-phosvitin and b-phosvitin with molecular weight of 160 000 and 190 000 Da, respectively, migrate in glycin-2,6-lutidine buffer without sodium dodecylsulfate (SDS) using electrophoresis. Under denaturating conditions with 0.5% SDS and Tris–glycine buffer, a-phosvitin and b-phosvitin dissociate into polypeptides with a molecular weight of 37.5, 42.5, 45 and 45 kDa, respectively (Abe et al., 1982). They also differ by their amino acid compositions, phosphorus contents, concentrations in carbohydrates, and solubility in presence of CaCl2 (Itoh et al., 1983). Phosvitin is relatively heat stable. Precipitation of phosvitin solution does not occur after heating at 100 °C between pH 4 to 7 (Mecham and Olcott, 1949). Itoh et al. (1983) observed no change in the electropherogram of aand b-phosvitin heated at 110 °C but the phosvitin bands were completely diffused at 140 °C.
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Fractionation of egg proteins and peptides 601 Lipovitellins (HDLs) In native egg yolks, lipovitellins or HDLs are linked to phosvitin to form the granular unit through phosphocalcic bridges. HDLs are in form of dimer of 400 kDa and are made up of 75–80% proteins and 20–25% lipids. HDLs can be separated by ion-exchange chromatography into a-HDL and b-HDL. Although they have similar chemical compositions (Bernardi and Cook, 1960), notable differences were observed in sialic acid and in their protein-bound phosphorus content (0.39 and 0.19% P, respectively) (Kurisaki et al., 1981). Each monomer of HDL consists of five major polypeptides with molecular weight ranging from 32 to 105 kDa, this latter being the main one. Plasma proteins in egg yolk Lipovitellenin (LDL) LDL is the main constituent of yolk (2/3 of the total yolk dry matter) and is mainly found in plasma but a residual portion is included in granules. The LDL structure was described as spherical particles with a lipid core surrounded by a layer of phospholipids and proteins (Evans et al., 1973). LDL is composed of 11–17% proteins and 83–89% lipids. There is little agreement among the various molecular weight values of the apoproteins of LDL published to date. Polypeptides reported range between 15 and 180 kDa (Anton et al., 2003; Yamauchi et al., 1976) and higher molecular weights up to 225–240 kDa have also been found (Itoh et al., 1986; Mine, 1998). Owing to its low density (0.98), LDL is soluble in aqueous solution whatever the pH and ionic conditions. Livetins Livetins comprise the water-soluble globular protein fraction which is composed of a-, b-, and g-livetins in the ratio 2:5:3, respectively, in the yolk (Bernardi and Cook, 1960). The molecular weights of a-, b-, g-livetins are 80, 45 and 150 kDa, respectively (Martin et al., 1957). The a- and b-livetins are thermolabile, whereas the g-livetins are more thermostable.
19.3 Biological activities of egg proteins and peptides and applications in the nutraceutical industry Biologically-active peptides can be produced in vitro through enzymatic hydrolysis of egg proteins. Most bioactive peptides are produced using gastrointestinal enzymes, such as pepsin, chymotrypsin and trypsin. For example, angiotensin-converting enzyme (ACE) inhibitory peptides from ovalbumin are usually produced by peptic hydrolysis. In general, they contain 2–20 amino acid residues per molecule but some may consist of more than 20 amino acids. These peptides can exhibit various activities including ACE-inhibitory, antihypertensive, antimicrobial, vasorelaxant and
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602 Separation, extraction and concentration processes immunomodulant actions. Some bioactive peptides can exert more than one activity (Table 19.4). 19.3.1 ACE inhibitor, antihypertensive and vasorelaxing activities Various peptides derived from proteolytic digestion of egg white proteins contain a number of peptide sequences having in vitro ACE-inhibitory activity. In vivo studies showed some antihypertensive and/or vasorelaxing activities in spontaneously hypertensive (SHR) rats models. Fujita et al. (1995) reported the first bioactive peptide, named ovokinin (OVA 358–365), derived from peptic digestion of ovalbumin. Ovokinin showed relaxing activity in canine mesenteric artery and possessed a high ACE-inhibitory activity, IC50 = 3.2 mM (Miguel et al., 2004). A hexapeptide corresponding to the 2–7 fragment of ovokinin (RADHPF) was isolated from a chymotryptic digest of ovalbumin. Ovokinin (2–7) showed lower ACEinhibitory activity (IC50 > 400 mM) but exerted a potent vasorelaxant and antihypertensive effect upon oral administration to SHR rats at a dosage of 10 mg kg–1 (Matoba et al., 1999). Non-hydrolyzed whole egg white did not have ACE-inhibitory property whereas its digestion with pepsin resulted in high ACE-inhibitory activity with IC50 = 55.3 mg mL–1 after hydrolysis for 3 h. Treatment with trypsin and chymotrypsin induced very poor ACE-inhibitory activity values, indicating the importance of the enzyme specificity in producing peptides with ACEinhibitory activities (Miguel et al., 2004). Furthermore, these workers showed that the active peptides could be enriched from the hydrolysate using ultrafiltration with a 3 kDa molecularweight-cut-off (MWCO) membrane. The permeate generated contained peptides with molecular mass <3 kDa that presented 10 times more ACE-inhibitory activity than the retentate (IC50 = 34.5 mg mL–1 and IC50 = 298.4 mg mL–1, respectively) and were identified as RADHPFL and YAEERYPIL. These two peptides exhibited a significant antihypertensive effect in SHR rats at a dose of around 2 mg kg–1 (Miguel et al., 2005). Proteolytic degradation with pepsin and pancreatic extract of RADHPFL and YAEERYPIL released two main fragments RADHP and YPI which decreased blood pressure at doses of 2 mg kg–1 after 2 h of administration but did not exhibit ACE-inhibitory activity (Miguel et al., 2006). Vasorelaxing activity has also been found in the sequences RADHP and YPI and show common features with other vascular relaxing peptides. Peptides with antihypertensive activity have also been generated by enzymatic hydrolysis of egg yolk. Oligopeptides of 1 kDa and less showed ACE-inhibitory and suppressed the development of hypertension in SHR rats in a dose-dependent manner after oral administration for 12 weeks (Yoshii et al., 2001).
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Table 19.4 Bioactive peptides derived from egg proteins Bioactive peptide
Fragment
Sequence
Enzyme
Bioactivity
Ovalbumin Ovalbumin Ovalbumin Ovalbumin Ovalbumin Lysozyme Lysozyme Ovalbumin Ovalbumin Ovalbumin Ovalbumin Ovalbumin Ovalbumin Ovalbumin Ovalbumin Lysozyme Phosvitin
Ovokinin Ovokinin (2–7) Unnamed Unnamed Unnamed Unnamed Unnamed Unnamed Unnamed Unnamed Unnamed Unnamed Unnamed Unnamed Unnamed Unnamed PPP
358–365 359–364 359–365 106–114 36–40 98–108 15–21 36–40 36–42 111–119 143–148 159–165 127–138 155–159 268–276 98–112 Unknown
FRADHPFL RADHPF RADHPFL YAEERYPIL SALAM IVSDGDGMNAW HGLDNYR SALAM SALAMVY YPILPEYLQ ELINSW NVLQPSS AEERYPILPEYL GIIRN TSSNVMEER IVSDGNGMNAWVAWR Unknown
Pepsin Chymotrypsin Pepsin Pepsin Pepsin Pepsin and trypsin Pepsin and trypsin Trypsin Trypsin Trypsin Trypsin Trypsin Chymotrypsin Chymotrypsin Chymotrypsin Clostripain Trypsin
Vasorelaxing/antihypertensive Vasorelaxing/antihypertensive ACE/antihypertensive ACE/antihypertensive ACE/antimicrobial Antimicrobial Antimicrobial Antimicrobial Antimicrobial Antimicrobial Antimicrobial Antimicrobial Antimicrobial Antimicrobial Antimicrobial Antimicrobial Antioxidant/anti-inflammatory
PPP, phosphopeptides.
Fractionation of egg proteins and peptides 603
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Protein precursor
604 Separation, extraction and concentration processes 19.3.2 Antimicrobial activity Lysozyme is an enzyme found in many cells, tissues and secretion of organisms (Jollès and Jollès, 1984). Lysozyme extracted from hen’s egg white exerts bacteriolytic activity on bacterial cells walls. It is most effective against some specific Gram-positive bacteria and, to a lesser extent, against Gramnegative bacteria owing to their composition of peptidoglycans (Pellegrini et al., 1992). Various strategies have been used to increase antimicrobial activity of lysozyme against Gram-negative bacteria. A number of chemical modifications aimed at inserting a hydrophobic moiety (Ibrahim et al., 1991; 1993), adding hydrophobic peptides to the C-terminus of lysozyme (Ibrahim et al., 1992; 1994) or lypophilizing of lysozyme by a fatty acid having a different chain length (Ibrahim et al., 1993; Liu et al., 2000). Nakamura et al. (1991, 1996) and Nakamura and Kato (2000) enhanced antimicrobial activity against Gram-negative bacteria by conjugating lysozyme with polysaccharides. Proteolytic digestion of lysozyme with clostripain has been reported to yield an antimicrobial peptide fragment (f98–112) (Ibrahim et al., 2001; Pellegrini et al., 1997). This peptide fragment has 15 amino acids and is located close to the C-terminal end of lysozyme; it possesses antibacterial activity against both Gram-positive and Gram-negative bacteria (Pellegrini et al., 1997). From a structural point of view, this sequence is a part of a helix-loop-helix domain located at the upper lip of the active site cleft of the lysozyme, which occurs in several bactericidal peptides (Ibrahim et al., 2001; Pellegrini, 2003). Chemical substitution of some amino acids to confer a net positive charge of the peptide 98–112 allowed further enhancement of these bactericidal properties. Hydrophobicity was also identified as an important factor in promoting interactions with bacterial membranes and had a strong impact on the bactericidal activity of this pentadecapeptide (Ibrahim et al., 2001; Pellegrini, 2003). Mine et al. (2004) characterized a novel antimicrobial peptide from chicken egg white lysozyme obtained by peptic and tryptic digestion. The peptide (98–108), located in the helixloop-helix domain and the peptide (15–21) possesses antimicrobial activity against Escherichia coli and Staphylococcus aureus. Commercial peptide mixtures from hen egg lysozyme, produced by partial enzymatic hydrolysis of lysozyme with pepsin, also exhibit a powerful antimicrobial activity to control Bacillus species in food products (Abdou et al., 2007). Ovalbumin contains bactericidal peptides that can be released by enzymatic digestion. Five antimicrobial peptides (SALAM, SALAMVY, YPILPEYLQ, ELINSW and NVLQPSS) from tryptic digests and peptides AEERYPILPEYL, GIIRN and TSSNVMEER from chymotryptic digests were found to be strongly active against Bacillus subtilus and to a lesser extent against the Gram-positive and Gram-negative bacteria studied (Pellegrini et al., 2004). Ovotransferrin displays in vivo and in vitro antibacterial action owing to both its iron-binding activity and its 92 residues N-terminal domain (Ibrahim et al., 1998; 2000). The cationic peptide of hen ovotransferrin, called OTAP
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Fractionation of egg proteins and peptides 605 92, corresponds to the sequence of ovotransferrin located at the lip of the iron-binding cleft of the N lobe of ovotransferrin. This peptide exhibits bactericidal activity against Gram-positive and Gram-negative bacteria by penetrating the bacterial outer membrane by self-promoted uptake and causing damage to the cytoplasmic membranes (Ibrahim et al., 2000). 19.3.3 Antioxidant activity Numerous phosphoserine residues grouped in the amino acids sequence of phosvitin, are responsible for the iron-binding capacity and they do confer to phosvitin a potential antioxidant activity. The metal-binding capacity of phosvitin can control iron-catalyzed or copper-catalyzed lipid oxidation to a Fe2+: phosvitin ratio of 30:1 and up to a Cu2+: phosvitin ratio of 1:1. Moreover, pasteurization does not affect the iron-binding capacity and antioxidant potential of phosvitin (Lu and Baker, 1986). Nakamura et al. (1998) produced a novel macromolecular antioxidant by conjugating phosvitin with galactomannan which could withstand a sterilization treatment at 121 °C for 15 min. Peptides generated from the digestion of various proteins are reported to have antioxidant activities. Jiang and Mine (2000) have developed phosvitin phosphopeptides (PPP) with molecular weight values of 1–3 kDa from tryptic hydrolysis following partial alkaline dephosphorylation. These peptides showed antioxidant activity against oxidative stress in human intestinal epithelial cells in an in vitro assay using Caco-2 cells (Katayama et al., 2006). In another study, PPP can up-regulate cellular glutathione biosynthesis associated enzymes activity and antioxidant activities in oxidative stress induced intestinal epithelial cells (Katayama et al., 2007). Although PPP have not yet been characterized, the antioxidant activity of PPP may be associated with molecular weight, phosphorus content, hydrophobicity and/ or amino acid composition. Ovalbumin was found to possess a strong antioxidant activity against linolenic acid and docosahexaenoic acid (Nara et al., 1995). Three short peptides from ovalbumin protein hydrolysates have been identified and it has been suggested that metal chelation plays an important role in their antioxidant activity (Hamachio and Hasegawa, 1989).
19.4 Available technologies for the fractionation of egg proteins and peptides, and applications in the nutraceutical industry Egg proteins are becoming important in the poultry product industry because of their technological and functional properties. Moreover, hen egg possesses many biologically active proteins that may have numerous applications in the food and pharmaceutical industries (Mine, 2007). Accordingly, there
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606 Separation, extraction and concentration processes is a need to develop efficient, simple and cost-effective methodologies for isolation and purification of egg proteins and their peptides. Figure 19.1 summarizes the technological options available for fractionating egg proteins and their peptides. The contemporary industrial processes comprise three main approaches alone or in combination, namely precipitation, chromatography and membrane processes. The appropriate selection of a given technological approach has to take into account some key physicochemical characteristics of the proteins/peptides to be separated from complex mixtures. Table 19.5 summarizes some of the structural properties of bioactive peptides derived from egg protein. These are often related to the amino acid sequence of the peptides. For example, ACE-inhibitory peptides commonly contain a positive side-chain charge on the C-terminal residue. This feature can be exploited by using ion-exchange chromatography and/ or membrane separation with charged polymeric material. Isoelectric precipitation Precipitation
Salting-out Organic solvents Gel permeation
Chromatography
Ion-exchange Reverse-phase/hydrophobic Microfiltration
Membrane separations
Ultrafiltration Nanofiltration
Fig. 19.1 Technological approaches and processes to fractionate egg proteins/ peptides. Table 19.5 Common structural properties of bioactive peptides Activity
Structural element
Antimicrobial
Short sequences (<40 residues) Helix-loop-helix motif Positive charge and hydrophobic properties Rich in SerP High amounts of His and hydrophobic amino acids Molecular size Positive charge in C-terminal residue Presence of aromatic and/or hydrophobic amino acids in C-terminal residue Arg and Tyr residues in N-terminal position
Antioxidant ACE inhibitory Vasorelaxing
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Fractionation of egg proteins and peptides 607 19.4.1 Precipitation Precipitation of egg proteins can be achieved by changes in the solvent properties including pH and/or ionic strength modifications and by addition of organic solvents. The most commonly used precipitating agents are ammonium sulfate, acetone or ethanol. Precipitation by ionic strength and/or pH modifications Purification of ovalbumin, the main egg-white protein, involves its precipitation at specific salt concentration, pH and temperature values. Ammonium sulfate or sodium sulfate are used for ovalbumin precipitation. Lysozyme is the main egg-white protein extracted on an industrial scale for commercial applications. Lysozyme can be precipitated from egg-white proteins by increasing the pH to 9.5 and by addition of sodium chloride to a final concentration of 5% (w/v) (Alderton and Fevold, 1946). Ovotransferrin has been purified with ammonium sulfate in crystalline form, both as the iron complex and as the iron free protein (Warner and Weber, 1951). Ovomucin can be precipitated by simple dilution with two or three volumes of water and by lowering pH to its pI of 4.5–5.0. A washing step with salt solution is required to dissociate ovomucin from lysozyme and finally salts are washed out with water (Cotterill and Winter, 1955). Purification of the immunoglobulin Y (IgY) involves removal of lipids and lipoproteins from egg yolk. Various strategies involving detergents such as SDS (Sriram and Ygeeswaran, 1999) and polysaccharides (Hatta et al., 1988; 1990), solvents (Sriram and Ygeeswaran, 1999) and polyethyleneglycol (Akita and Nakai, 1993) were used to remove lipids from egg yolk. Salt precipitation can thereafter be achieved for IgY purification. Ammonium sulfate and sodium sulfate were used and concentration levels were dependent on the IgY yield and purity (Akita and Nakai, 1992). Although purification by precipitation of proteins for preparative purposes can be achieved, these methods suffer from major deficiencies. Precipitation by salts leads to a protein extract and to by-products having a high salt concentration. Several solubilizations and crystallization cycles must be performed to obtain highly purified proteins. Precipitation by organic solvents Addition of an organic solvent to protein solution results in important modifications of the dielectric constant of the medium and to the weakening of protein interactions with water. Organic solvents have an affinity for the hydrophobic surfaces of the proteins and this result in denaturation of the proteins along with precipitation. However, low temperatures (–5 to 0 °C) are necessary to minimize protein denaturation. Ovomucoid is the main egg white protein isolated by using organic solvent. First, precipitation of other egg white proteins is eliminated with 2.7% trichloroacetic acid at pH 3.5. Then, ovomucoid is precipitated by addition of acetone (Lineweaver and Murray, 1947).
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608 Separation, extraction and concentration processes After salt precipitation, using solvent such as cryoethanol gave an IgY isolate of 93% purity (Akita and Nakai, 1992); ethanol is used in large-scale IgG production (Horikoshi et al., 1993). Precipitation by organic solvents is widely used in the chemical industry. However, the necessity of several precipitation steps to obtain a purified protein make this process extremely difficult to scale up and the low temperature restricts the scale of production. Moreover, the use of organic solvent is considered undesirable in the food and nutraceutical industrial environment. 19.4.2 Chromatographic methods Egg proteins are characterized by specific physicochemical characteristics such as their size, pI, amino acid composition, ability to bind metal ions (ovotransferrin and lysozyme), protein–protein interaction potential (lysozyme and ovomucin) or ligand-binding capacity (avidin and flavoprotein). Based on these characteristics, several methods, including gel-permeation, ionexchange and reversed-phase chromatography, were used for the separation or purification of egg white and yolk proteins. Gel-permeation chromatography Gel-permeation chromatography (GPC) separates molecules according to the difference in their size. This technique is based on the penetration of molecules into the cavities of a macroporous support, mostly made from hydrophilic gels of dextran, agarose or polyacrylamide. In general, molecules with a hydrodynamic diameter smaller than the diameter of the pores in the support diffuse into the matrix, whereas molecules with larger diameters are excluded, thus passing through more quickly. Several proteins of egg-white and yolk proteins have been isolated by GPC. Purification of ovomucin from egg-white proteins was carried out using a Sepharose 4B and Superose 6 preparative-grade columns. Gel permeation on Superose 6 permitted simultaneous purification of ovomucin and lysozyme, but with Sepharose 4B another compound eluted with the ovomucin and this may be ovostatin (Awade et al., 1994; Young and Gardner, 1972). GPC was used in combination with other chromatographic procedures. After gel permeation on Superose 6, ovotransferrin and ovalbumin were isolated by anion-exchange chromatography on Q Sepharose Fast Flow. The two-step purification procedure gave 80, 100, 80 and 100% purity for ovomucin, lysozyme, ovotransferrin and ovalbumin, respectively. In another study, a better separation of ovomucin and lysozyme was obtained with a Superose 12 HR 10/30 column and anion-exchange or reversed-phase high-performance liquid chromatography (RP-HPLC) were proposed to purify other egg white proteins (Awade and Efstathiou, 1999). GPC was also used to isolate phosvitin, the very low-density lipoproteins from delipidated egg yolk proteins (Tsutsui and Obara, 1984). Abe et al.
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Fractionation of egg proteins and peptides 609 (1982) separated phosvitin into a- and b-phosvitin by gel permeation on a Sephadex G-200 column. These two components differ in their amino acid composition, their carbohydrate content and their precipitation in the presence of calcium (Itoh et al., 1983). GPC is most useful in the separation of egg proteins owing to their wide range of molecular mass distribution. However, GPC allows compounds having a molecular size between two relatively close thresholds to be separated. Because of the fragile nature of many soft gels used for GPC and because interaction with proteins reduces the potential utilization of GPC for egg protein fractionation, large-scale processes may not be easily applicable. Ion exchange chromatography Ion-exchange chromatography (IEC) is the most widely used large-scale method for the purification of proteins and other charged molecules. In cation-exchange chromatography, positively charged molecules are attracted to a negatively charged solid support. Conversely, in anion-exchange chromatography, negatively charged molecules are attracted to a positively charged solid support. A variety of resins may be used to fractionate lysozyme from egg white proteins. Among weakly acidic cation-exchange resins, Duolite C-464 exhibited an efficient method for separation and recovery of active lysozyme (86%). However, a significant amount of avidin co-eluted with lysozyme (Durance and Nakai, 1988; Li-Chan et al., 1986). Roy et al. (2003) described an integrated easily scalable process that is used to simultaneously purify the major egg white proteins. This technique involves a cation exchanger, Streamline™ SP and differential precipitation, followed by dye–ligand chromatography. Purified proteins were obtained and the yields of lysozyme, ovomucoid and ovalbumin were 77, 94 and 98%, respectively. Guérin-Dubiard et al. (2005) developed a procedure for fractionating the whole egg white using three successive chromatography steps. High-purity levels of lysozyme (95%), ovotransferrin (89%), ovalbumin (91%) and flavoprotein (100%) were obtained. Connelly and Taborsky (1961) separated phosvitin by use of stepwise salt elution on a DEAE cellulose column into two sub-fractions: a major fraction called 0.30, and a minor one called 0.35 phosvitin. The fractions had approximately the same amino acid composition but differed in their metal content and chemical stability at alkaline pH. Recently, Castellani et al. (2003) described a new purification method that includes a first extraction step based on insolubility of Mg2+/phosvitin salts and a second step by ion-exchange chromatographic fractionation avoiding organic solvents. Purification gives a- and b-phosvitin that are free from contaminants, highly purified (>98%) and metal-free phosvitin. IEC is suitable for fractionating egg proteins with high purity and, until now, the application of IEC on an industrial scale concerned the major egg white proteins.
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610 Separation, extraction and concentration processes Reversed-phase high-performance liquid chromatography The separation mechanism in RP-HPLC is based on the hydrophobic interaction of proteins/peptides with a non-polar stationary phase. Itoh et al. (1991) investigated the recovery of hydrophobic egg white proteins using RP-HPLC at room temperature. Under conventional gradient elution conditions, the recovery of ovalbumin was poor. However, under fast separation conditions, the recovery of the proteins was dramatically improved. In contrast, under both chromatographic conditions approximately 100% of lysozyme was recovered. For RP-HPLC of lipid-free egg yolk proteins solvents containing formic acid were used. Although several peaks were unidentified, three groups (granules proteins, low-density lipoprotein apoproteins, and livetins) were separated (Sheumack and Burley, 1988). RP-HPLC is not used to fractionate egg white and yolk proteins as much as other chromatographic methods. A gradient of a mixture of a non-polar solvent is required for elution and protein denaturation may occur during RP-HPLC process. 19.4.3 Membrane processes Pressure-driven membrane-based separation processes comprise microfiltration (MF), ultrafiltration (UF), nanofiltration (NF) and reverse osmosis (RO) and are used for protein separation/purification (Cheryan, 1998). Depending on MWCO, various membranes can be used for the separation of different molecular weight proteins. MF membranes are suitable for the separation of particles in the size range of 0.1–10 mm whereas UF membranes are used for macromolecules with a molecule weight of 1–300 kDa. Compared with UF, NF membranes have a smaller pore size and retain smaller organic molecules, 200–2000 Da. RO drives water molecules from a low to a high concentration region and results in the concentration of salt molecules on the retentate side of the membrane. Size is the main sieving mechanism in MF and RO whereas the separation mechanism is normally explained in terms of charge and/or size effects in UF and NF. UF is mainly used for protein concentration, desalting, clarification and protein fractionation. Protein fractionation is strongly influenced by operating and physicochemical parameters such as pH and ionic strength that affect the protein–protein and protein–membrane interactions and thus the selectivity. Model egg white protein solutions were fractionated with modified and unmodified 50 kDa MWCO polysulfone UF membranes (Eshani et al., 1997). Electrostatic exclusion prevented ovotransferrin and lysozyme from permeating the membrane and this led to a permeate containing almost pure ovalbumin. Fractionation was obtained at pH 4.8 without salts on both UF membranes. In another study, UF polysulfone membrane pre-treated with myoglobin enhanced lysozyme purification from chicken egg white compared with that obtained with the native membrane (Ghosh and Cui, 2000a). Ghosh and Cui (2000b) reported the effect of pH on fractionation of
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Fractionation of egg proteins and peptides 611 lysozyme from chicken egg white. A high transmission of lysozyme was observed with increasing pH showing the influence of pH on the purification of lysozyme. In some instances, a combination of UF membranes is useful for protein fractionation. A two-step ultrafiltration process using 50 and 25 kDa MWCO membranes was achieved to purify lysozyme. In a first step, lysozyme was enriched in the UF permeate of 50 kDa MWCO. Then, the 25 kDa MWCO UF membrane was used to purify the lysozyme. This combination gave a high productivity and purity for the separation process. Similarly, a two-stage UF technique was used to separate ovalbumin from chicken egg white. In the first stage, ovalbumin was retained by a 30 kDa flat disk polyethersulfone membrane; this retentate was fractionated using 50 kDa flat disk polyethersulfone membrane in the second stage. A high-purity ovalbumin (98.7%) could be produced using a two-step UF process under the experimental conditions studied (Datta et al., 2009). Separation of lysozyme from chicken egg white has been investigated by using 30 kDa hollow fiber polysulfone membranes. Higher lysozyme transmission (>90%) was observed under optimized conditions and moderately pure lysozyme (80–90%) could be obtained by carrying out a diafiltration step (Ghosh et al., 2000). Production of ACE-inhibitory peptides from egg white proteins with various proteolytic enzymes was carried out using a membrane reactor. Egg white proteins hydrolyzed with thermolysin produced the highest ACE-inhibitory potential (IC50 = 54.1 mg mL–1). Thermolysin hydrolysate was fractionated using UF membrane of MWCO of 10, 3 and 1 kDa. A lower IC50 (17.2 mg mL–1) was obtained in the UF permeate with 1 kDa MWCO (Chiang et al., 2008). By a modification of polymeric forms of lysozyme under optimal conditions of UF, a preparation comprising 53.3% of lysozyme polymeric forms, i.e. 33.2% of dimer and 20.1% of trimer, was obtained. Modified lysozyme preparation by UF showed the highest bacteriostatic activity against Gramnegative bacteria selected (Lesnierowski et al., 2009). A water-soluble plasma protein from egg yolk granules was obtained with various simple water dilutions, followed by centrifugation or filtration. Two factors were critical, pH and egg yolk dilution, for IgY recovery. Optimum recovery of IgY (93–96%) was obtained by a six-fold water dilution at pH 5.0–5.2 with incubation for 6 h at 4 °C. A 100 kDa MWCO extracted IgY from water-soluble plasma upon the separation of crude IgY by ammonium sulfate precipitation. A maximal recovery of IgY (> 98%) was obtained and this value was similar to the degree of purity obtained by gel filtration (Akita and Nakai, 1992). A serial filtration approach, involving dilutions, paper filtration and delipidation using hydrophobic filters or using different UF membranes, was developed by Kim and Nakai (1996; 1998). The delipidated water-soluble fraction was thereafter purified using 100 kDa UF membranes and both high recoveries (72–89%) and purity (74–99%) were obtained. Although the fractionation of proteins through a membrane depends
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612 Separation, extraction and concentration processes on physicochemical parameters such as pH and salt concentration (Kim and Nakai, 1996; 1998), the surface properties of the membrane are also important for the successful protein fractionation. Hernandez-Campos et al. (2010) purified IgY from egg yolk water-soluble protein by UF–diafiltration with different membranes. The best selectivity and purification factors were obtained without salt at pH 5.7 and 6.7 using polyethersulfone and modified polyethersulfone of 100 MWCO membrane, respectively. Compared with chromatographic methods, membrane separation techniques offer the advantage of lower cost and ease of scale-up for commercial production. Membrane-based separation techniques are powerful tools for fractionation/purification of proteins and bioactive peptides. UF membranes can be used as preliminary step for the removal of enzymes and non-hydrolyzed proteins, and to further fractionate the peptide mixture. Various combinations of precipitation and UF using selective membranes can be applied for the fractionation of food proteins. In addition to selective membranes, chromatographic methods can be employed successfully to isolate specific bioactive compounds from a complex mixture. UF membranes were successfully used to enrich peptide fractions. Membranes consisting of negatively charged materials such as NF were used to desalt or to fractionate acidic peptides from hydrolysate mixture (Chay Pak Ting et al., 2007; Wijers et al., 1998). 19.4.4 Other separation methods A new isolation method was described for IgY. Use of food-grade products such as sodium alginate or l-carrageenan was investigated for precipitation of yolk lipoproteins from the water-soluble fraction to purify IgY. Polysaccharides exhibit satisfactory yields of IgY (Hatta et al., 1988; 1990). In another study, the optimal separation of IgY was achieved with 0.15% pectin and at pH 5.0, implying that the interactions between polysaccharides and lipoproteins were mainly ionic bonds, hydrophobic interactions and hydrogen bonds (Chang et al., 2000).
19.5 Conclusion and perspectives Considering that the discovery of novel bioactive peptides and their possible functions and health benefits are constantly increasing, there is a growing need to develop new technologies for the production of specific bioactive peptides from hydrolysate mixtures. Membrane-based separation techniques have traditionally been used to separate molecules of different sizes and lowmolecular-weight components from proteins. UF, alone and in combination with other techniques such as NF and chromatography has shown its potential to separate peptide mixtures on a large scale. However, the further research into membrane areas (such as membrane materials and chemistry, © Woodhead Publishing Limited, 2010
Fractionation of egg proteins and peptides 613 module configurations) for the complete separation/purification of proteins/ peptides is still needed. The addition of affinity ligands to chelate proteins/ peptides could be useful in obtaining maximal selectivity by combining with membrane processes.
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614 Separation, extraction and concentration processes Connelly C and Taborsky G (1961), Chromatographic fractionation of phosvitin, J Biol Chem, 236, 1364–1368. Cotterill O J and Winter A R (1955), Egg white lysozyme. 3. The effect of pH on the lysozyme-ovomucin interaction, Poult Sci, 34, 679–686. Cunningham F E and Lineweaver H (1965), Stabilization of egg white proteins to pasteurization temperature above 60 °C, Food Technol, 19, 136–141. Datta D, Bhattacharjee S, Nath A, Das R, Bhattacharjee C and Datta S (2009), Separation of ovalbumin from chicken egg white using two-stage ultrafiltration technique, 66, 353–361. Donovan J W and Mapes C J (1976), A differential scanning calorimetric study of conversion of ovalbumin to S-ovalbumin in eggs, J Sci Food Agric, 27, 197–204. Durance T D and Nakai S (1988), Simultaneous isolation of avidin and lysozyme from egg albumen, J Food Sci, 53, 1096–1102. Eshani N, Parkkinen S and Nyström M (1997), Fractionation of natural and model eggwhite protein solutions with modified and unmodified polysulfone UF membranes, J Membr Sci, 123, 105–119. Evans R J, Bauer D H, Bandemer S L, Vaghefi S B and Flegal C J (1973), Structure of egg yolk very low density lipoprotein. Polydispersity of the very low density lipoprotein and the role of lipovitellenin in the structure. Arch Biochem Biophys, 154, 493–500. Fujita H, Usui H, Kurahashi K and Yoshikawa M (1995), Isolation and characterization of ovokinin a bradykinin B1 agonist peptide derived from ovalbumin, Peptides, 16, 785–790. Ghosh R and Cui Z F (2000a), Protein purification by ultrafiltration with pre-treated membrane, J Membr Sci, 167, 47–53. Ghosh R and Cui Z F (2000b), Purification of lysozyme using ultrafiltration, Biotechnol Bioeng, 68, 191–203. Ghosh R, Silva S S and Cui Z F (2000), Lysozyme separation by hollow fibre ultrafiltration, Biochem Eng J, 6, 19–24. Greengard O, Sentenac A and Mendelsohn N (1964), Phosvitin, the iron carrier of egg yolk, Biochim Biophys Acta, 90, 406–407. Guérin-Dubiard C, Pasco M, Hietanen A, Quiros del Bosque A, Nau F and Croguennec T (2005), Hen egg white fractionation by ion-exchange chromatography, J Chromatogr A, 1090, 58–67. Hamachio Y and Hasegawa M (1989), EPA powder as a functional ingredient, New Food Ind, 31, 12–16. Hatta H, Kim M and Yamamoto T (1990), A novel isolation method for hen egg yolk antibody, ‘IgY’, Agric Biol Chem, 54, 2531–2535. Hatta H, Sim J S and Nakai S (1988), Separation of phospholipids from egg yolk and recovery of water-soluble proteins, J Food Sci, 53, 425–427, 431. Hernandez-Campos F J, Brito-De La Fuente E and Torrestiana-Sanchez B (2010), Purification of egg yolk immunoglobulin (IgY) by ultrafiltration: effect of pH, ionic strength and membrane properties, J Agric Food Chem, 58(1), 187–193. Horikoshi T, Hiraoka J, Saito M and Hamada S (1993), IgG antibody from hen egg yolks: purification by ethanol fractionation, J Food Sci, 58, 739–742. Ibrahim H R, Iwamori E, Sugimoto Y and Aoki T (1998), Identification of a distinct antibacterial domain within the N-lobe of ovotransferrin, Biochim Biophys Acta, 1401, 289–303. Ibrahim H R, Kato A and Kobayashi K (1991), Antimicrobial effects of lysozyme against Gram-negative bacteria due to covalent binding of palmitic acid, J Agric Food Chem, 39, 2077–2082. Ibrahim H R, Kobayashi K and Kato A (1993), Length of hydrocarbon chain and antimicrobial action to Gram-negative bacteria of fatty acylated lysozyme, J Agric Food Chem, 41, 1164–1168.
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Fractionation of egg proteins and peptides 615 Ibrahim H R, Sugimoto Y and Aoki T (2000), Ovotransferrin antimicrobial peptide (OTAP 92) kills bacteria through a membrane damage mechanism, Biochim Biophys Acta, 1523, 196–205. Ibrahim H R, Thomas U and Pellegrini A (2001), A helix-loop-helix peptide at the upper lip of the active site cleft of lysozyme confers potent antimicrobial activity with membrane permeabilization action, J Biol Chem, 276, 43767–43774. Ibrahim H R, Yamada M, Kobayashi K and Kato A (1992), Bactericidal action of lysozyme against Gram-negative bacteria due to insertion of a hydrophobic pentapeptide into its C-terminus, Biosci Biotechnol Biochem, 56, 1361–1363. Ibrahim H R, Yamada M, Matsushita M, Kobayashi K and Kato A (1994), Enhanced bactericidal action of lysozyme to Escherichia coli by inserting a hydrophobic pentapeptide into its C-terminus, J Biol Chem, 18, 5059–5063. Ishihara H, Takahasi N, Ito J, Takeuchi E and Tejima S (1981), Either high-manonose-type or hybrid-type oligosaccharide is linked to the same asparagine residue in ovalbumin, Biochim Biophys Acta, 669, 216–221. Itoh T, Abe Y and Adachi S (1983), Comparative studies on the a- and b-phosvitin from hen’s egg yolk, J Food Sci, 48, 1755–1757. Itoh T, Kubo M and Adashi S (1986), Isolation and characterization of major apoproteins from hen’s egg yolk granule, J Food Sci, 51, 1115–1117. Itoh T, Miyakazi J, Sugawara H and Adachi S (1987), Studies on the characterization of ovomucin and chalaza of the hen’s egg, J Food Sci, 52, 1518–1521. Itoh H, Nimura N, Kinoshita T, Nagae N and Nomura M (1991), Fast protein separation by reversed-phase high-performance liquid chromatography on octadecylsilyl-bonded nonporous silica gel. II. Improvement in recovery of hydrophobic proteins, Anal Biochem, 199, 7–10. Jiang B and Mine Y (2000), Preparation of novel functional oligophosphopeptides from hen egg yolk phosvitin, J Agric Food Chem, 48, 990–994. Jollès P and Jollès J (1984), What’s new in lysozyme research?, Mol Cell Biochem, 63, 165–189. Katayama S, Ishikawa S I, Fan M Z and Mine Y (2007), Oligophosphopeptides derived from hen egg yolk phosvitin up regulate g-glutamylcysteine synthetase and antioxidant enzymes against oxidative stress in Caco-2 cells, J Agric Food Chem, 55, 2829–2835. Katayama S, Xu X, Fan M Z and Mine Y (2006), Antioxidant stress activity of oligophosphopeptides derived from hen egg yolk phosvitin in Caco-2 cells, J Agric Food Chem, 54, 773–778. Kato I, Schrode J, Kohr W J and Laskowski Jr M (1987), Chicken ovomucoid: determination of its amino acid sequence, determination of trypsin reactive site, and preparation of all three of its domains, Biochemistry, 26, 193–201. Kim H and Nakai S (1996), Immunoglobulins separation from egg yolk: a serial filtration system, J Food Sci, 61, 510–513. Kim H and Nakai S (1998), Simple separation of immunoglobulin from egg yolk by ultrafiltration, J Food Sci, 63, 485–490. Kurisaki J K, Yamauchi H, Ishiki H and Ogiwara S (1981), Differences between a- and b-lipovitellin from hen egg yolk, Agric Biol Chem, 45, 699–704. Lesnierowski G, Kijowski J and Cegielska-Radziejewska R (2009), Ultrafiltrationmodified chicken egg white lysozyme and its antibacterial action, Int J Food Sci Technol, 44, 305–311. Li-Chan E, Nakai S, Sim J, Bragg D B and Lo K V (1986), Lysozyme separation from egg white by cation exchange column chromatography, J Food Sci, 51, 1032–1036. Li-Chan E and Nakai S (1989), Biochemical basis for the properties of egg white, Crit Rev Poult Biol, 2, 21–57. Lineweaver H and Murray C W (1947), Identification of the trypsin inhibitor of egg white with ovomucoid, J Biol Chem, 171, 565–581.
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616 Separation, extraction and concentration processes Liu S T, Sugimoto T, Azakami H and Kato A (2000), Lipophilization of lysozyme by short and middle chain fatty acids, J Agric Food Chem, 48, 265–269. Losso J N, Bogumil R and Nakai S (1993), Comparative studies of phosvitin from chicken and salmon egg yolk, Comp Biochem Physiol, 106, 919–923. Lu C L and Baker R (1986), Characteristics of egg yolk phosvitin as an antioxidant for inhibiting metal-catalyzed phospholipid oxidations, Poult Sci, 65, 2065– 2070. Martin W G, Vandegaer J E and Cook W H (1957), Fractionation of livetin and the molecular weights of the a- and b-components, Can J Biochem Physiol, 35, 241–250. Matoba N, Usui H, Fujita H and Yoshikawa M (1999), A novel anti-hypertensive peptide derived from ovalbumin induces nitric oxide-mediated vasorelaxation in an isolated SHR mesenteric artery, FEBS Lett, 452, 181–184. Matsuda T, Watanabe K and Nakamura R (1982), Immunochemical studies on thermal denaturation of ovomucoid, Biochim Biophys Acta, 707, 121–128. Mecham D K and Olcott H S (1949), Phosvitin, the principal phosphoprotein of egg yolk, J Am Chem Soc, 71, 3670–3679. Miguel M, Aleixandre A, Ramos M and Lopez-Fandino R (2006), Effect of simulated gastrointestinal digestion on the antihypertensive properties of ACE-inhibitory peptides derived from ovalbumin, J Agric Food Chem, 54, 726–731. Miguel M, Lopez-Fandino R, Ramos M and Aleixandre A (2005), Short-term effect of egg-white hydrolysate products on the arterial blood pressure of hypertensive rats, Br J Nutr, 94, 731–737. Miguel M, Recio I, Gomez-Ruiz J A, Ramos M and Lopez-Fandino R (2004), Angiotensin I-converting enzyme inhibitory activity of peptides derived from egg white proteins by enzymatic hydrolysis, J Food Prot, 67, 1914–1920. Mine Y (1998), Adsorption behaviour of egg yolk low-density lipoproteins in oil-in-water emulsions, J Agric Food Chem, 46, 36–41. Mine Y (2007), Egg proteins and peptides in human health-chemistry, bioactivity and production, Curr Pharm Des, 13, 875–884. Mine Y, Ma F and Lauriau S (2004), Antimicrobial peptides released by enzymatic hydrolysis of hen egg white lysozyme, J Agric Food Chem, 52, 1088–1094. Mine Y, Noutomi T and Haga N (1990), Thermal induced changes in egg white proteins, J Agric Food Chem, 38, 2122–2125. Nakamura R and Ishimaru M (1981), Changes in the shape and surface hydrophobicity of ovalbumin during its transformation to S-ovalbumin, Agric Biol Chem, 45, 2775–2780. Nakamura S, Gohya Y, Losso J N, Nakai S and Kato A (1996), Protective effect of lysozyme-galactomannan or lysozyme-palmitic acid conjugates against Edwardsiella tarda infection in carp, Cyprinus carpio, FEBS Lett, 383, 251–254. Nakamura S, Kato A and Kobayashi K (1991), New antimicrobial characteristics of lysozyme-dextran conjugate, J Agric Food Chem, 39, 647–650. Nakamura S and Kato A (2000), Multi-functional biopolymer prepared by covalent attachment of galactomannan to egg white proteins through naturally occurring Maillard reaction, Nahrung, 44, 201–206. Nakamura S, Ogawa M, Nakai S, Kato A and Kitts D D (1998), Antioxidant activity of a Maillard-type phosvitin–galactomannan conjugate with emulsifying properties and heat stability, J Agric Food Chem, 46, 3958–3963. Nara E, Miyashita K and Ota T (1995), Oxidative stability of PC containing linoleate and docosahexaenoate in an aqueous solution with or without chicken egg albumin, Biosci Biotech Biochem, 59, 2319–2320. Pellegrini A (2003), Antimicrobial peptides from food proteins, Curr Pharm Des, 9, 1225–1238. Pellegrini A, Hülsmeier A J, Hunziker P, and Thomas U (2004), Proteolytic fragments of ovalbumin display antimicrobial activity, Biochim Biophys Acta, 1672, 76–85.
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Fractionation of egg proteins and peptides 617 Pellegrini A, Thomas U, Bramaz N, Klauser S, Hunziker P and von Fellenberg R (1997), Identification and isolation of bactericidal domain in chicken egg white lysozyme, J Appl Microbiol, 82, 372–378. Pellegrini A, Thomas U, von Fellenberg R and Wild P (1992), Bactericidal activity of lysozyme and aprotinin against Gram-negative and Gram-positive bacteria related to their basic character, J Appl Bacteriol, 72, 180–187. Perlman G E (1952), Enzymatic dephosphorylation of ovalbumin and plakalbumin, J Gen Physiol, 25, 711–726. Powrie W D and Nakai S (1985), Characteristics of edible fluids of animal origin: eggs, in Fennema O, Food Chemistry 2nd edition, New York, Marcel Dekker Inc., 829–855. Radomski M W and Cook W H (1964), Fractionation and dissociation of the avian lipovitellins and their interaction with phosvitin, Can J Biochem, 42, 395–406. Roy I, Rao M V S and Gupta M N (2003), An integrated process for purification of lysozyme, ovalbumin and ovomucoid from hen egg white, Appl Biochem Biotechnol, 111, 55–63. Sheumack D D and Burley R W (1988), Separation of lipid-free egg yolk proteins by high-pressure liquid chromatography using solvents containing formic acid, Anal Biochem, 174, 548–551. Sriram V and Ygeeswaran G (1999), Improved recovery of immunoglobulin fraction from egg yolk of chicken immunized with AsialoGM1, Russ J Immunol, 4, 131–140. Stadelman W J and Coterril O J (1973), Egg science and technology, Westport: Avi Publishing, p. 314. Stein P E, Leslie A G, Finch J T, Turnell W J, McLaughlin P J and Carrell R W (1990), Crystal structure of ovalbumin as a model for the reactive centre of serpins, Nature, 347, 99–102. Stein P E, Leslie A G, Finch J T and Carrell R W (1991), Crystal structure of uncleaved ovalbumin at 1.95 Å resolution, J Mol Biol, 221, 941–959. Tsutsui T and Obara T (1984), Preparation and characterization of phosvitin from hen’s egg yolk granule, Agric Biol Chem, 48, 1153–1160. Van Het Schip F D, Salmallo J, Broos J, Ophuis J, Mojet M, Gruber M and Geert A B (1987), Nucleotide sequence of a chicken vitellogenin gene, J Mol Biol, 196, 245–260. Warner R C and Weber I (1951), The preparation of crystalline conalbumin, J Biol Chem, 191, 173–180. Watanabe K, Matsuda T and Sato Y (1981), The secondary structure of ovomucoid and its domain as studied by circular dichroism, Biochim Biophys Acta, 667, 242–250. Watanabe K, Shimoyamada M, Onizuka T, Akiyama H, Niwa M, Ido T and Tsuge Y (2004), Amino acid sequence of a-ovomucin in hen egg white ovomucin deduced from cloned cDNA, DNA Seq, 15, 251–261. Wijers M C, Pouliot Y, Gauthier S F, Pouliot M and Nadeau L (1998), Use of nanofiltration membranes for the desalting of peptide fractions from whey protein enzymatic hydrolysates, Le Lait, 78, 621–632. Williams J, Elleman T C, Kingston I B, Wilkins A G and Kuhn K A (1982), The primary structure of hen ovotransferrin, Eur J Biochem, 122, 297–303. Wiseman R L, Fothergill J E and Fothergill L A (1972), Replacement of asparagines by aspartic acid in hen ovalbumin and a difference in immunochemical reactivity, Biochem J, 127, 775–780. Yamauchi K, Kurizaki J and Sasago K (1976), Polypeptide composition of hen’s egg yolk very low-density lipoprotein, Agric Biol Chem, 40, 1581–1586. Yoshii H, Tachi N, Ohba R, Sakamura O, Takemaya H and Itani T (2001), Antihypertensive effect of ACE inhibitory oligopeptides from chicken egg yolks, Comp Biochem Physiol C Toxicol Pharmacol, 128, 27–33.
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618 Separation, extraction and concentration processes Young L L and Gardner F A (1972), Preparation of egg white ovomucin by gel filtration, J Food Sci, 37, 8–11. Young A C, Tilton R F and Dewan J C (1994), Thermal expansion of hen egg-white lysozyme. Comparison of the 19 Å resolution structures of the tetragonal form of the enzyme at 100 K and 298 K, J Mol Biol, 235, 302–317.
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Supercritical-fluid extraction of lycopene from tomatoes 619
20 Supercritical-fluid extraction of lycopene from tomatoes J. Shi and S. Jun Xue, Agriculture and Agri-Food Canada, Canada, Y. Jiang, The Chinese Academy of Sciences, China and X. Ye, Zhejiang University, China
Abstract: Several process parameters for supercritical CO2 fluid extraction – such as pressure, temperature, flow rates, co-solvent or modifier concentrations, resident time, moisture content, particle sizes, and particle size distribution – have individual or combined effects on the recovery of lycopene from tomatoes. The solubility and bioactivity of lycopene, and the composition of the extract and extraction yield can be affected. Improved processing conditions and reduced cost are required if the extraction of lycopene from tomato materials using the supercritical CO 2 fluid extraction process is to become more economical at low throughputs. Key words: bioactivity, lycopene, supercritical fluid extraction, tomato.
20.1 Introduction Lycopene is a phytochemical responsible for the red pigments found in plants. It is a non-provitamin A carotenoid that plays an important role in the biosynthesis of many carotenoids. Recently, the extraction of lycopene and other carotenoids has attracted attention owing to their biological and physiochemical properties, particularly those which possess natural antioxidant activities. Antioxidants have been associated with disease prevention and reduction (Negre-Salvayre et al., 2006; Papas, 1999; Wu, et al., 1999). Structurally, lycopene is an acyclic, open-chain, C40 polyisoprenoid unsaturated carotenoid having 13 double bonds, of which 11 are conjugated, arranged in a linear array, and it has a molecular formula of C40H56 (Fig. 20.1). Because of this unique molecular structure, lycopene has high antioxidant activity and singlet oxygen quenching ability that is thought to be beneficial
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620 Separation, extraction and concentration processes
Fig. 20.1 Structure of lycopene.
to human health. The antioxidant protective effect of lycopene and tomatoes has been shown in vitro as well as in vivo (Breinholt et al., 2000; Hadley et al., 2003; Rao and Agarwal, 1998; Suganuma and Inakuma, 1999; Willcox et al., 2003). Therefore, foods containing lycopene are of considerable interest. The singlet oxygen quenching ability of lycopene has been found to be three times greater than that of b-carotene and a-tocopherol (Stahl and Sies, 1992, 1996). Epidemiological studies and investigations have shown important roles of lycopene and other carotenoids in free radical inactivation and fat peroxidation inhibition (Krinsky and Rock, 1998). The health benefits of lycopene-rich diets and its effect on minimizing the risk of cardiovascular ailments and various forms of cancers have been reported (Palozza, 1998; Olson, 1986). More recent studies have demonstrated the important functions of lycopene and other carotenoids in the pathophysiology of chronic diseases (Rao and Agarwal, 1999; Rao and Rao, 2007). Increasingly, in vivo and in vitro clinical studies have shown its protective effect against the growth of tumor cells and its ability to protect against cardiovascular (Arab and Steck, 2000) and coronary heart diseases, and cancer (Clinton, 1998). With its 11 conjugated and two non-conjugated double bonds, lycopene is a more efficient antioxidant (singlet oxygen quencher) than b-carotene, a-carotene, and a-tocopherol (Mascio et al., 1989; Shi, 2002; Shi and Le Maguer, 2000). As the most abundant carotenoid in tomatoes, lycopene is the dominant pigment responsible for the color of tomatoes. Lycopene occupies the largest portion, about 80–90%, of the carotenoids in ripe tomatoes, followed by b-carotene, phytoene, and the other minor carotenoids. Concentrations of lycopene in tomatoes approach 50–100 mg kg–1, and it has a higher colour intensity than b-carotene (Shi and Le Maguer, 2000). Other sources of lycopene include pink grapefruit, guava, watermelon, autumn olive, and apricots. Ripe tomato skins contain approximately five times more lycopene than the pulp. The industrial processing of tomato products produces wastes such as seeds and skins. However, reuse of those wastes is limited to animal feed. Even though ripe tomatoes are the most abundant source of lycopene, over 90% is located in the skin, which constitutes the greater part of the waste, and is a potential natural source for lycopene extraction (Shi and Le Maguer, 2000). Lycopene content in plant materials is dependent on the species and the temperature at which growth and maturation occurs. Factors such as climate trends significantly alter the amount of lycopene in the material, thus affecting the yield. Lycopene is the last pigment to appear during maturation and its
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Supercritical-fluid extraction of lycopene from tomatoes 621 development is inhibited by temperatures higher than 30–32 °C, whereas other carotenoids degrade at these temperatures. As a result, the pigment concentrations in plant materials may vary between crop cycles (Shi et al., 1999b). The concentration of lycopene also changes during ripening and storage. Liu et al. (2009) reported that the accumulated lycopene contents of tomatoes did not change significantly during the first four days of storage. However, between days 4 and 21, the lycopene content increased. A growing interest in and demand for healthy, environmentally safe, and cost-efficient products has driven the research and application of new technologies in the food, pharmaceutical and cosmetic industries. Extraction processes are commonly used to enrich and detoxify food through the removal of targeted components from natural products. Organic solvents are used in conventional methods for the extraction of bio-compounds such as lycopene from plant materials. However, these solvents not only generate environmentally hazardous problems and requests for expensive disposal procedures for the chemical extraction solvents, but also chemical residues remaining in the final products become a major safety concern. With increasing government restrictions reflecting consumer concerns on food safety, alternative and reliable extraction techniques are of great interest. Extraction processes are needed for ‘additive-free natural’ products. It is clear that the current concern for safety in food products has increased interest in ‘green’ extraction techniques, instead of the conventional organic solvent extraction processes. The determination of lycopene in food products and development of a safe ‘green’ extraction process to complement the fortification of functional foods is of great public interest. One possible environmentally friendly alternative is supercritical fluid extraction (SFE), in particular using supercritical CO2 ‘green technology’ because it is physiologically harmless, environmentally safe, non-explosive, exhibits high selectivity as a result of low viscosity, high diffusivity, and liquid-like density, as well as being readily available and easily removed from products (Simandi et al., 2002). Because of the CO2 supercritical state (31 °C and 7.38 MPa), procedures should allow supercritical operation of thermally labile compounds that would be easily degraded at high temperatures. SFE has attracted growing interest for the recovery of natural compounds for large-scale industrial production over recent decades (O’Day and Rosenau, 1982). SFE is successfully and widely used for the extraction of lycopene from ripe tomatoes (Cadoni et al., 2000) and tomato processing wastes (Baysal et al., 2000; Kassama et al. 2008; Ollanketo et al., 2001; Rozzi et al., 2002; Sabio et al., 2003; Topal et al., 2006; Vasapollo et al., 2004). One significant thermodynamic advantage of using supercritical fluid is its ease of separation from the extracted solutes by simply modifying the operation conditions, either pressure or temperature. Supercritical fluids have liquidlike densities that give superior mass transfer characteristics compared with organic solvents, and are characterized as low-viscosity and high-diffusivity fluids. In addition, supercritical fluids have low surface tension leading to
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622 Separation, extraction and concentration processes easy penetration into the porous biological matrix of plant material while releasing the targeted components.
20.2 Supercritical-fluid extraction (SFE) of lycopene Supercritical fluids are, by definition, at a temperature and pressure greater than or equal to the critical temperature and pressure of the fluid. They actually have physical properties somewhere between those of a liquid and a gas. Supercritical fluids are able to spread out along a surface more easily than a true liquid because they have lower surface tensions than liquids. The supercritical CO2 is fed at high pressure by means of a pump, which pressurizes the extraction tank and also circulates the supercritical fluid throughout the system. Figure 20.2 demonstrates an example of a typical single-stage supercritical CO2 extraction system. Once the supercritical CO2 reaches the equilibrium state in the extraction vessel, the extraction process proceeds with the manipulation of pressure and temperature to achieve the ideal operating conditions. The mobile phase, consisting of the supercritical CO2 and the solubilized lycopene component, are transferred to the separator where the solvating power of the fluid is decreased by increasing the temperature and or decreasing the pressure of the system. The lycopene extract precipitates in the separator whereas the supercritical CO2 is either released or recycled back to the extractor. As the solution containing the extracts of lycopene compounds leave the extractor and flow to the first separation vessel via the pressure regulator, a paste of lycopene oleoresins settles to the bottom as they separate and this is collected, whereas the remaining solution goes to the second-stage separator where the fractionation of the lycopene components occurs. A third stage of separation might be required for the complete isolation of pure lycopene
Back pressure valve
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Fig. 20.2 Schematic diagram of a typical single-stage supercritical fluid extraction system with CO2 (modified from Shi et al., 2006).
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Supercritical-fluid extraction of lycopene from tomatoes 623 components. In commercial-scale processing, multiple extraction vessels are normally used to enhance operation efficiency and throughput (Fig. 20.3). The process requires an intermittent batch system because it is interrupted at the end of the extraction period, when the pressure must be released so the extraction vessel can be switched out of the extractor loop. By having two or more interchangeable extraction vessels, the unloading and reloading of a vessel can occur while extraction of a previously charged vessel is in progress. This reduces the downtime and improves the overall production efficiency.
20.3 Factors affecting lycopene yield Many studies have successfully extracted lycopene from tomato materials by SFE. The technology has been continuously developed and users have established environmentally sound and complimentary safe techniques for extracting food-grade, bioactive components used as natural ingredients from agricultural materials. There have been numerous studies on the effects of various independent variables – such as pressure, temperature, flow rates, and co-solvent or modifier concentrations – on lycopene yield using SFE (Baysal et al., 2000; Ollanketo et al., 2001; Ciurlia et al., 2009; Kassama et al., 2008; Yi et al., 2009, Shi et al., 2009a; 2009b; Vaughn et al., 2008). Two-stage separation vessels
Storage tank
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Fig. 20.3 Schematic diagram of commercial scale multi-stage supercritical fluid extraction system used to fractionate bioactive components.
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624 Separation, extraction and concentration processes The optimization of pressure, temperature, and flow rate is necessary to obtain a high recovery of lycopene. Baysal et al. (2000) observed the highest recovery (54%) of lycopene obtained at 55 °C, 30 MPa, and 4 kg h–1 with 5% ethanol as modifier. Ollanketo et al. (2001) recovered 94% of lycopene in 15 min at 40 MPa and 110 °C. Rozzi et al. (2002) obtained a maximum recovery (61%) at 80 °C, 35 MPa, and 2.5 mL min–1. Yi et al. (2009) studied the effects of SFE parameters on the recovery of lycopene and antioxidant activity. In this study, the maximum total lycopene content of 31.25 mg g–1 of raw tomato was extracted at the highest temperature of 100 °C, 40 MPa and 1.5 mL min–1. Process parameters such as solvent flow rate, residence time, moisture content, particle sizes, and particle size distribution in conjunction with supercritical pressures and temperatures are key parameters for achieving optimum results. Most of these parameters can have individual or combined effects on the extraction rate, bioactivity, and composition of the extract. For example, the residence time can have an important influence on the composition of the extracted compound. 20.3.1 Effects of temperature, pressure and flow rate on recovery of lycopene Effect of pressure Plate IIa (between pages 292 and 293) shows that pressure significantly influences the rate of extraction. Higher density causes a double effect that increases the solvation power and reduces the interaction between the fluids and matrix, resulting in a decrease in the diffusion coefficient. Excessive pressure also increases the compactness of the sample matrix, thus reducing the pore sizes and apparently reducing the mass transport which eventually diminishes the yield (Tonthubthimthong et al., 2001). It is commonly considered that an increase in pressure results in an increase in CO2 density, increasing the solvating power of the supercritical fluid. Thus, higher pressure is responsible for quantitative recoveries and stronger interactions between the fluid and the matrix (Shi et al., 2009a; Topal et al., 2006). Effect of temperature Temperature is a parameter with significant influence on SFE yield; thus manipulating it could have an adverse implication on the process and yield. Plate IIb shows the general trend of increased extraction yield as temperature increases relative to the pressure. The extraction yields increased both with temperature and pressure, whereas the effect of temperature is more significant than that of pressure. Vági et al. (2007) maximized the recovery of lycopene at 46 MPa and 80 °C by SFE. At 40–100 °C, maximum yields were obtained for extraction of lycopene. The highest temperature gave the maximum lycopene content (Rozzi et al., 2002). Cadoni et al. (2000) reported that at 85 °C a maximum of 65% of lycopene was recovered from the pulp
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Supercritical-fluid extraction of lycopene from tomatoes 625 and skins of ripe tomatoes. However, with further elevation of temperature, degradation and isomerization of lycopene was often encountered (Shi et al., 1999a, 2009b, 2002; Shi and Le Maguer, 2000; Yi et al., 2009). It is well-known that increasing the temperature reduces the solvent density and consequently decreases the recovery of lycopene at constant pressure. At the same time, higher temperature promotes the solubility of lycopene and increases the recovery by increasing mass transfer of lycopene in the matrix, and from the matrix to the CO2 extraction medium (Marsili and Callahan, 1993; Shi et al., 2007d). Therefore, the increase in temperature could have either a positive or a negative effect, as a result of the balance between CO2 density (r) and solute vapour pressure. The increase in the recovery of lycopene by SFE depends more on the solute’s vapour effect. Although lycopene is not stable during long heating times and was approximately 53.5% degraded after 60 min at 100 °C through isomerization or auto-oxidation (Boskovic, 1979; Mayeaux et al., 2006; Shi and Le Maguer, 2000; Shi et al., 1999a). However, in the thermal processing of tomato products, the total lycopene contents still increased in the products owing to the elevated release of lycopene from the tomato tissue matrix (Dewanto et al., 2002; Seybold et al., 2004; Toor and Savage, 2006). For SFE, the enhanced release of lycopene from the insoluble fibre portion of tomato skin may also contribute to the increased recovery. Effect of CO2 flow rate In an extraction, the flow rate of supercritical CO2 has an influence on the extraction efficiency. The flow rate controls the amount of solvent (e.g. CO2) to be used and the extraction time. Extraction is a diffusion-based process, with the solvent required to diffuse into the matrix, and the extracted component to diffuse out of the matrix into the solvent at equilibrium state. The flow rate also greatly affects solubility of extracted components in fluid and equilibrium of diffusion between solvent and extracted components during extraction. If the concentration of extracted components is much higher than the solvent, the solubility of extracted components in the solvent could be limited and slow down the extraction process. Therefore, optimization of flow rate is important for extraction time and cost efficiency. For example, to maximize the rate of extraction, the flow rate should be high enough for the extraction to be completely diffusion limited, but this is very wasteful of solvent. In contrast, the lower flow rate would minimize the amount of solvent used, but the extraction should be completely solubility limited, and thus would take a very long time. Therefore, the optimum flow rate is probably somewhere in the region where both solubility and diffusion are significant factors. At optimum temperature and pressure, the extraction process is performed under conditions in which the CO2 flow rate is variable. The linear correlation is observed for the flow rates and the extraction rates. The extraction rate increased with increasing flow rate. When the intra-particle diffusion
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626 Separation, extraction and concentration processes resistance is considered as a factor in the material matrix, more CO2 is required to make contact with the targeted components. Thus, the extraction rate, is a function of CO2 consumed: the greater the amount of supercritical gas used, the higher the yield obtained (Friedrich and Pryde, 1984). When the temperature, pressure, and overall extraction time are kept constant, the effect of various solvent flow rates on the carotenoid yield is minor. Plate IIc shows the effect of flow rate on total lycopene yield at constant pressure (30 MPa) or temperature (70 °C). An increase in flow rate from 1.0 to 2.0 mL min–1 did not show a significant change in the yield of lycopene (Yi et al., 2009). Results indicated that the higher flow rate could produce faster extractions and higher recoveries since the higher flow rate could overcome the interface resistance for lycopene transport from the tomato skin tissue to the CO2 fluid (Hawthorne et al., 1995). However, excessively high flow rates would produce undesirable results owing to a decrease in contact times between the solvent and the solute (Topal et al., 2006). This may force the solvent to leave the system without reaching its solubility limit (Anitescu et al., 1997). High flow rates may also cause the sample to compact and restrict the fluid flow, thus reducing the amount of solvent that comes in contact with the solute (Franca and Meireles, 2000). Increasing the CO2 flow rate would increase fluid diffusivity and improve the heat and mass transfer within the system, because the resulting fluid temperature approaches the wall temperature more quickly (Shi et al., 2007a). The effect of solvent flow rate on mass transfer rate and total yield in the constant extraction rate period is significant. Effect of operating time Residence time is an important factor that influences yield and the economic viability of the process. When optimizing extraction conditions, the extraction time of the process is an important parameter. For the extraction of lycopene by SFE Baysal et al. (2000) found that the highest carotene yield was obtained at an extraction time of 2 h, as opposed to 1 or 3 h. A 1-h extraction may not suffice for the maximum amount of carotenoids to be dissolved in the solvent, whereas over 3 h there is an increase in the occurrence of degradation. The duration and intensity of thermal processing are directly correlated to the degree of isomerization and degradation of the targeted compounds (Baysal et al., 2000). Therefore, to minimize the degradation effect, it is advisable to reduce the extraction time as much as possible. In addition, optimizing extraction conditions towards the shortest extraction time possible with the maximum extraction recovery is ideal from the point of view of the extract quality and processing cost. Effect of moisture content of raw materials Water represents approximately 80–90% of the total weight of fresh plant materials and interferes with the effectiveness of the SFE (Hopper and King, 1991). Thus, the initial moisture content may improve or impede the
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Supercritical-fluid extraction of lycopene from tomatoes 627 yield results depending on the polarity of the compound. The non-polar compounds such as lycopene produce smaller yields with higher moisture content (Anitescu et al, 1997; Chao et al., 1993). Water has a small but finite solubility in supercritical CO2, therefore, it can also be extracted with the targeted components. The presence of a little water (less than 10%) aided in the extraction of lycopene, particularly at the beginning of the process (Shi et al., 2006). A moisture content of 7–18% had negligible effect on extractability of oil. However, above an 18% moisture content, the presence of water affected the extraction and the efficiency decreased (Nagy and Simándi, 2008). The high water content in the matrix inhibited the flow of SCF by changing surface tension and contact angles as a result of phase interaction. Therefore, water removal in most cases frees the internal pores and thus increases the mass transport intensity. Higher moisture contents cause a higher probability of formation of a thin film of water between the matrix and the SCF phase. Possible effects of water include an increase in distance that carotenoids must travel to reach the solvent and swelling of the cell matrix, which have negative and positive effects on the process, respectively. Effect of particle size Particle size is a factor that influences the extraction recovery, having a significant impact on the flow behaviour of SCF in the sample matrix. The extraction process and diffusivity of the extraction medium are hindered by large particle sizes (>550 mm) because of the distance which the lycopene travels from the inside of the particle to the surface (Shi et al., 2007a). By reducing the size of the particles, components such as carotenoids have less distance to travel. Thus, the collective amount of surface area with which the solvent is in contact is increased, an important factor during the initial stages of extraction (Ollanketo et al., 2001). The particle size is affected by sample pre-treatment. For example, the methods of drying (air, oven, vacuum, or freeze drying), milling, and other mechanical, ultrasonic, high-pulsed electronic field, and non-mechanical treatments can change the particle size by attrition or size reduction. Numerous experimental results have indicated that smaller particle sizes are ideal for extraction processes, and the manner in which the solid material is crushed and powdered also affect the results (Ferreira and Meireles, 2002; Spanos et al., 1993). The degree to which the material is crushed determines the amount of cells that are broken, and this has a positive correlation with recovery. The smaller the particle size, the larger is the surface area, and the bioactive components are released more easily. However, if the particle size is exceedingly small that can favour channelling which commonly results in a decrease in recovery (Ziémons et al., 2005). Only 40% lycopene recovery was obtained from powdered tomato skins crushed with a household blender, relative to the amount achieved when sea sand was used to assist the grinding procedure (Ollanketo et al., 2001). Therefore, optimizing particle
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628 Separation, extraction and concentration processes size significantly improves the yield of extraction. Sun and Temelli (2006) noted that the optimal range of particle size was 0.25–0.5 mm for lutein, b-carotene, and a-carotene. However, for certain processes, it is important to be cautious of the degree of crushing. Gouveia et al. (2006) studied the effects of slight, moderate and complete crushing, which gave 20%, 40%, and 65% carotenoid recovery, respectively, and also found that samples that have been extremely crushed may not be affected by the addition of entrainers.
20.4 Effects of pressure and temperature on the antioxidant activity of lycopene Recovery is increased with increasing temperature owing to increasing solubility of the targeted compounds; therefore, for economic reasons, food industries prefer to increase the temperature to achieve higher yields. However, the bioactivity of the lycopene may be compromised by heat because of the instability of lycopene at higher temperatures (Mayer-Miebach et al., 2005; Shi et al., 2002b, 2004b, 2008, 2009b; Vega et al., 1996; Yi et al., 2009). Yi et al. (2009) reported that the recovery and bioactivity of lycopene increased when the temperature increased from 20 to 60 °C. The recovery of lycopene increased slowly and its antioxidant activity declined above 80 °C. Pol et al. (2004) made a remarkable observation that lycopene recovery was increased with increasing temperature, but the bioactivity of the extracted lycopene did not follow the trend of the recovery of lycopene extracts at high temperature. No significant effects of pressure or flow rate on the antioxidant activity were observed (Yi et al., 2009). Cadoni et al. (2000) and Shi and Zhou (2006) reported that the results obtained from SFE of lycopene from tomatoes are poorly reproducible because of the decomposition and isomerization of lycopene during the extraction process. Yi et al. (2009) investigated the effects of the two major supercritical fluid extraction parameters (temperature and pressure) on the antioxidant activities of the extracts as shown in Fig. 20.4. The recovery increased with the increase in pressure, whereas antioxidant activity decreased slightly. The recovery of total lycopene increased with temperature, but the antioxidant activity of the lycopene-rich extract decreased with increasing temperature (Hackett et al., 2004; Yi et al., 2009). At 25 and 50 °C, lycopene in an oleoresin is degraded predominately through oxidation, whereas isomerization increases at 75 and 100 °C. If lycopene is isomerized, the proportion of lycopene isomers is changed, that then further influences the antioxidant activity. In general, the change in antioxidant activity in lycopene-rich extracts is mainly influenced by their lycopene content, proportions of the isomers of lycopene, and the synergistic effects of carotenoids in the extract (Shi et al., 2004a,b, 2004c, 2007c; Shixian et al., 2005). The total antioxidant activity of
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0.20
20 0.15
15
0.10
10 5 20
25
30 Pressure (MPa) (b)
35
Torlox of lycopene (mM mg–1)
35
Torlox of lycopene (mM mg–1)
Lycopene concentration (mg g–1)
Supercritical-fluid extraction of lycopene from tomatoes 629
0.05 40
Fig. 20.4 Changes of lycopene yield (mg –1g raw material) (d) and antioxidant activity (mM mg–1 Torlox) of the extract (m) (a) at 40 to 100 °C, 30 MPa and flow rate of 1.5 mL min–1, and (b) at 20–40 MPa, 70 °C and flow rate of 1.5 mL min–1 (modified from Yi et al., 2009).
the extract, that is the antioxidant activity of each unit of lycopene multiplied by the total lycopene amount in the extract, firstly significantly increases as the temperature increases, and then slightly decreases as temperature is kept at a high temperature. The antioxidant activity of extracts first increased with increasing lycopene contents in the extracts between 40 and 80 °C, and then started to decline thereafter. The slight decrease above 90 °C might be affected by the degradation of other carotenoids (Cocero et al., 2003). Dewanto et al. (2002) found that thermal processing increased the total antioxidant activity of tomato from 4.13±0.36 to 6.70±0.25 mmol g–1 of vitamin C equivalents after 30 min of heating at 88 °C. The total antioxidant activity probably increases because larger amounts of lycopene are released from the matrix during thermal processing. HPLC chromatograms of the extracts
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630 Separation, extraction and concentration processes obtained at 50 and 90 °C are shown in Fig. 20.5. All-trans lycopene is the major configuration eluting with a retention time of 15.87 min under both conditions. Three major cis-isomer peaks are detected around 6.8, 8.9, and 16.1 min of elution time in extracts prepared at 50 °C. When the extraction temperature is elevated to 90 °C, total lycopene increased 1.18 fold, whereas all-trans lycopene contents increased only 1.03 fold and total cis-isomers increased more than 2.13 fold, respectively. Moreover, more than three major peaks of cis-isomers are detected (Fig. 20.5) because of new cis-isomer generations. The higher increase in cis-isomers than total and all-trans lycopene indicates the transformation of all-trans to cis-isomers. As extraction vessel temperature increased from 50 to 90 °C, the ratio of all-trans to total cislycopene isomers changed from 1.67 to 1.40 (Table 20.1). Further increases in the operating temperature resulted in dramatic isomerization. The lower ratio of trans- to cis-isomers obtained at high temperature indicated higher lycopene bioavailability owing to the cis-isomers generated (Boileau et al., 1999). The isomerization of lycopene in tomato samples during thermal processing above 75 °C, and the decrease in the trans- to cis-isomer ratios from 9:1 to 2:3 after heating at 140 °C, were demonstrated (Cocero et al.,
4 cis isomers mAU
3
trans-lycopene
2 5-cis
1 0
0
5
10
15 20 Time (min) (a)
25
30
25
30
trans-lycopene
4 mAU
cis isomers 3 2
5-cis
1 0
0
5
10
15 20 Time (min) (b)
Fig. 20.5 Chromatograph of SFE extract obtained at (a) 50 °C and (b) 90 °C at 30 MPa and flow rate of 1.5 mL min–1 (modified from Yi et al., 2009). mAu, milli absorbance unit. © Woodhead Publishing Limited, 2010
Supercritical-fluid extraction of lycopene from tomatoes 631 Table 20.1 All-trans and cis-isomers contents and ratio in the SFE extract obtained at 40–100 °C, 30 MPa and a flow rate of 1.5 mL min–1 (modified from Yi et al., 2009) Temperature Total content All-trans (mg g–1) content (mg g–1) 100 90 80 70 60 50 40
29.46 26.12 22.69 14.66 13.85 12.00 10.35
16.75 15.26 13.76 9.04 8.74 7.50 6.52
Total cis Ratio Ratio Ratio* content (trans:total) (cis:total) (trans:cis) (mg g–1) 12.71 10.86 8.93 5.62 5.11 4.50 3.83
0.57 0.58 0.61 0.62 0.63 0.63 0.63
0.43 0.42 0.39 0.38 0.37 0.37 0.37
1.32c 1.40c 1.54bc 1.62ab 1.71a 1.67a 1.70a
*The different letters shows the values that are significantly different (P <0.05) (n = 3).
2003; Mayer-Miebach et al., 2005; Shi et al., 2002a; 2002b; Xianquan et al., 2005).
20.5 Effect of co-solvent and modifiers in lycopene extraction The use of co-solvents or modifiers during SFE is key to enhancing the extraction efficiency and cost effectiveness of the processes. Joslin et al. (1996) indicated two significant attributes of co-solvents: the interaction between the co-solvents or modifiers with the solute (direct effect) and the co-solvents or modifiers with solvent interactions (indirect effect). Co-solvents or modifiers when used in small doses (1 to 5% mol) can change the overall characteristics of the extraction fluid in terms of polarity, solvent strength, and specific interactions. These changes, in turn, can significantly alter the density and compressibility of the SCF. Additionally, they can improve selectivity for desired components and facilitate selective fractional separations. The effect of a modifier depends upon the nature of the solute to be extracted (Walsh et al., 1987). The first basis for co-solvent or modifier selection is the increased solubility of the target in the modified CO2 fluid (Pourmortazavi and Hajimirsadeghi, 2007; Shi et al., 2007b, 2009a). Carbon dioxide has adequate solvent properties for extraction of targeted compounds. The addition of a small amount of a liquid modifier or co-solvent such as water, oil, or ethanol can significantly enhance the extraction of non-polar compounds such as carotenoids. Another advantage of selecting a co-solvent or modifier is the ability to distort and swell the matrix, favouring the penetration of the CO2 into the matrix in order to extract the analyte (Casas et al., 2007). Ethanol, water, and vegetable oil under various temperatures and pressures are often used. The addition of these co-solvents or modifiers significantly increases the total recovery because the solvent power of supercritical CO2 could be © Woodhead Publishing Limited, 2010
632 Separation, extraction and concentration processes increased by the addition of small amounts of co-solvents or modifiers. The effect of co-solvents or modifiers is dependent on their concentration in the supercritical phase, and this is determined by the phase behaviour of the mixture under operating conditions. It also should be noted that the addition of large amounts of modifier will change the critical parameters of the mixture (Pourmortazavi and Hajimirsadeghi, 2007). 20.5.1 Effect of ethanol as modifier Baysal et al. (2000) used ethanol at different concentrations (5, 10, and 15% w/w) to recover b-carotene and lycopene from tomato paste. The addition of ethanol increases the bulk density of supercritical CO 2 because of the higher density of the co-solvent and clustering of supercritical CO 2 molecules around the co-solvent (Güçlü-Üstündag and Temeli, 2005). The solubility of trans-b-carotene both in pure and ethanol-modified supercritical CO2 increases with temperature from 40 to 60 °C (Sovová et al., 2001). The ethanol modifier increased the solubility by one order of magnitude and the increase in solubility was proportional to the square root of the modifier concentration (Güçlü-Üstündag and Temeli, 2005; Shi et al., 2009a, 2009b; Sovová et al., 2001). Because of the high molar mass and elongated shape of the lycopene molecule, the modifier dilutes the extract, reducing the viscosity, and thereby enhancing the flow of the extract through the extractor. The increase in recovery showed a linear correlation with the increase in ethanol concentration at 45 °C. However, the increased recovery (6%) was lower when the ethanol concentration was increased from 10 to 15% than the increase (11%) when the ethanol concentration was increased from 5 to 10% (Shi et al., 2009b). The results suggest that selecting the proper concentration of modifier could improve extraction efficiency and reduce the operating cost. 20.5.2 Effect of water as a modifier Water enhances the analyte–modifier–matrix interaction because water opens the pores and causes the matrix to swell thereby allowing the supercritical fluid better access to the analytes to bring them out of the matrix. The recovery of lycopene increased significantly with an increase in water concentration in tomato skin materials (Yi et al., 2009). An excess of water (greater than 18%) causes mechanical difficulties such as the extraction fluid (e.g. CO2) clogging the restrictor (Casas et al., 2007) because the high water content makes material transport difficult. Previous studies reported that the extraction of lycopene from tomatoes varied in terms of the corresponding moisture content. Baysal et al. (2000) extracted lycopene from tomato paste waste having 24% moisture content, whereas Rozzi et al. (2002) extracted lycopene from tomato waste materials with a moisture content of 48.4%. As the moisture content of the tomato skins in the reported experiment was
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Supercritical-fluid extraction of lycopene from tomatoes 633 7%, the addition of 15% water would be approximately equal to a moisture content of 19% (Yi et al., 2009). Even though water is only about 0.3% soluble in supercritical CO2 (Lehotay, 1997), water has been used as a cosolvent in a number of SFE applications. In addition, water aids the extraction process by increasing the polarity of the supercritical fluids, thus enabling higher recoveries of relatively polar species (Pourmortazavi and Hajimirsadeghi, 2007). The adsorption of water onto the polar sites of the plant material matrix weakens the bond between the extractable component and the matrix, leading to an increased vapour pressure of the extract and enhanced rate of desorption (Brady et al., 1987). 20.5.3 Effect of edible oil as a modifier Considering the lipophilic properties of lycopene, edible oils used as alternative modifiers were proposed to enhance SFE (Krichnavaruk et al., 2008; Sun and Temelli, 2006; Vasapollo et al., 2004). One advantage is that such edible oil does not need to be subsequently separated from the product. Examples of the use of vegetable oil as a co-solvent for SFE include the extraction of lycopene from tomatoes using hazelnut oil (Vasapollo et al., 2004); the extraction of carotenoids from carrots using canola oil (Sun and Temelli, 2006), and the extraction of astaxanthin from Haematococcus pluvialis using vegetable oil (Krichnavaruk et al., 2008). Vasapollo et al. (2004) tested the effect of 10% hazelnut oil on extractable lycopene content from dried tomato (6% of moisture, average particle size of about 1 mm) under varying pressures and at two different temperatures (40 and 60 °C). Recovery of lycopene with olive oil as a modifier was higher than those obtained with ethanol as a modifier (Shi et al., 2009b). Olive oil, and most probably other vegetable oils, are good solvents for lycopene in the liquid state as they enhance the extraction yield by increasing the solubility of lycopene in the modified supercritical fluid. The slight reduction in recovery obtained with 15% olive oil as a modifier may be attributed to an excess of the oil plugging or clogging the pores of the support. 20.5.4 Effect of binary and ternary modifiers Very little information is available on the influence of the behaviour of two combined modifiers (ethanol and water, water and oil, and ethanol and oil) on the extraction of lycopene. Shi et al. (2009b) compared the effects of binary and ternary modifiers such as water, ethanol, and olive oil on the recovery of lycopene by SFE under various conditions. The results showed that the addition of various quantities of a mixture of ethanol and water resulted in higher lycopene contents than for the single modifier in equal amounts. The addition of a mixture of ethanol and olive oil showed a significant synergistic effect on the recovery compared with the addition of a mixture of ethanol and water, and the enhancement increased with concentration of both ethanol and
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634 Separation, extraction and concentration processes olive oil (Table 20.2). The highest recovery 73.3% was achieved by adding 10% ethanol and 10% olive oil. Olive oil has H-bonding ability, which could link to ethanol and contribute to the interaction between modified supercritical CO2 and the analyte. Owing to the fat-soluble characteristic, the addition of olive oil enhanced the solubility in supercritical CO2 modified with olive oil. Moreover, the addition of ethanol could remove more olive oil and lycopene. The recoveries obtained with a modifier mixture of water and oil were higher than that obtained by modifier mixture of water and ethanol, and lower than that obtained by a modifier mixture of oil and ethanol. The combined water and oil mixtures protect the lycopene extract from oxidative degradation (Boon et al., 2008). 20.5.5 Effects of modifiers on the ratio of all-trans to cis-isomers in lycopene extract According to Mayer-Miebach et al. (2005), there is not a dramatic effect on the isomerization of lycopene below 75 °C. In contrast, all-trans-lycopene was degraded by up to 25% when synthetic lycopene-containing water-in-oil emulsions were heated at temperatures even below 70 °C (Mayer-Miebach and Spiess, 2003). The proportions of cis-isomers in all the extracts obtained with olive oil as a modifier increased compared with extracts obtained with modifier of ethanol or water because the cis-isomers are less likely to crystallize and are more efficiently dissolved in lipophilic solutions such as oil (Shi et al., 2009b).
20.6 Solubility of lycopene in supercritical fluids Solubility refers to the amount of solute that will dissolve in a given amount of solution at thermodynamic equilibrium. Interactions between solvent and Table 20.2 Effects of modifiers on the ratios of all-trans to cis-lycopene in SFE from tomato skins at 350 MPa and 45 and 75 °C (E: ethanol; W: water; O: olive oil) (modified from Shi et al., 2009b) 45 °C
15%E 15% W 15% O 10% E + 10% W 10% E + 10% O 10% W + 10% O 5% E + 5% W + 5% O
75 °C
Recoveries (%)
Ratio (all-trans:cis)
Recoveries (%)
Ratio (all-trans:cis)
42.9 33.3 38.7 39.4 47.0 36.6 37.1
38.2 29:1 33:2 34:2 41:2 31:2 31:2
51.7 48.8 58.2 62.5 73.3 56.8 57.9
45:2 42:2 49:3 54:3 61:3 49:3 49:3
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Supercritical-fluid extraction of lycopene from tomatoes 635 solute molecules influence the tendency of a particular substance to dissolve in a solvent (Shi et al., 2007a). The solubility of targeted compounds is important when establishing the extraction parameters. To obtain the maximum recovery from a source material, the solubility of a component in the solvent should be as high as possible. It also dictates the amount of solvent necessary to optimize the extraction efficiency. The insolubility of lycopene in water causes a problem when attempting to extract it from natural materials. Although it is soluble in organic solvents such as benzene, chloroform and methylene chloride, these are toxic and therefore not ideal candidates for pharmaceutical, food, or cosmetic purposes. This is one of the main reasons why lycopene solubility in supercritical fluids has become of great interest (Cadoni et al., 2000). 20.6.1 Factors affecting the solubility of lycopene in SFE There are numerous factors which affect the solubility of lycopene in supercritical fluids. The solvating power of fluids in the supercritical phase is very sensitive to changes in temperature and pressure. Other parameters, including the presence of co-solvents and modifiers and compound morphology, also have to be considered. Pressure An increase in pressure at a constant temperature increases the solvating power of the fluid. This allows extraction of a wider variety of carotenoids, and thus results in an increase in solvent density, which can lead to a rapid increase in the solubility of the targeted component. Because density increases with pressure, solubility tends to increase with pressure. If the solvating power of the supercritical fluid is increased, it may contribute to a decrease in the diffusion coefficient at higher density, and the interaction between the fluid and the matrix may be reduced (Gordon and Bauernfeind, 1982). Johannsen and Brunner (1997) demonstrated that the solvating power of CO2 increases as the density increases, and found that the increase in solubility of b-carotene in supercritical CO2 occurs when the pressure of the system is increased. Corresponding to this trend, the amount of required CO2 is reduced as the extraction pressure increases. Sabio et al. (2003) noted that, after a certain point, the change in recovery with increasing pressure/density was insignificant. Subra et al. (1997) indicate that at lower pressures, the uncertainty of solubility measurements increases. However, low solubility of the condensed phase, particularly those of b-carotene, occurs at pressures close to the critical pressure of the solvent (Sakaki, 1992). The effect of pressure is enhanced by the addition of modifiers such as canola oil. The solubility of carotenoids increases as the pressure increases; however, modifiers enhance this effect under the same experimental conditions (Shi et al., 2009b; Sun and Temelli, 2006). A decrease in solubility from an increase in pressure at a constant
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636 Separation, extraction and concentration processes temperature is a phenomenon that has been observed in other experiments (Calvey and Page, 1990; Luque de Castro et al., 1994; Tsai et al., 2007). The density of CO2 rises when the pressure is increased; therefore, this may cause a reduction in the diffusivity of the solute through the solvent. Tonthubthimthong et al. (2001) suggested that an increase in pressure may cause the plant material to compress or to become more compact, thus restricting the solvent’s ability to enter the matrix. Pressure may affect the selectivity of the solvent for particular solutes, meaning that a particular range of pressures may favour the extraction of lycopene whereas other conditions are more suitable for other compounds. Figure 20.6 shows the relationship between the solubility of lycopene in supercritical CO2 and pressure at constant temperatures. This relationship suggests that the optimum solubility conditions are not at the highest combination of temperature and pressure tested. However, there is a significant relationship between temperature and lycopene solubility. Increases in temperature result in higher solvent volatility, and the solubility of lycopene is increased as a result. Although there were decreases in solvation power when the pressure was increased at a relatively high temperature, the magnitude of the change in solubility became greater with increasing temperature. In the study of the effects of temperature and pressure on the solubility of lycopene in supercritical CO2 fluid by Shi et al. (2009a), the results showed a considerable decrease in solubility from 25 to 30 MPa at 75 °C. However, the solubility is greater than that at 30 MPa and 65 °C. This may be because the density of the solvent is reduced, making it easier for the lycopene to diffuse into the solid. In addition, the magnitude of the solvent density changes is smaller at elevated pressures. The drop in
Solubility, y2 ¥ 106 (mol solute/106 mol mixture)
2.5
2
1.5
1
0.5
0 100
200
300 Pressure (bar)
400
500
Fig. 20.6 Solubility isotherms for lycopene: (®) 50 °C; () 60 °C; () 70 °C; (d) 80 °C. Lines represent trends and symbols represent experimental data points (modified from Shi et al., 2009a).
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Supercritical-fluid extraction of lycopene from tomatoes 637 solubility from 25 to 30 MPa at 65 and 75 °C may occur because the change in density of supercritical CO2 is not enough to compensate for any matrix packing of the source material that occurs (Shi et al. 2009a). The degree of the change in CO2 density as the temperature increases may be influenced by the pressure. However, the solubility at 75 °C is higher than that at 65 ºC at 30 MPa, despite a relatively severe drop in solubility when the pressure was raised, probably because the increased molecular interactions between the supercritical CO2 and lycopene eclipses the inhibiting effect caused by the increased solvent density. Temperature Elevations in temperature instigate changes in the density of supercritical CO2 and lycopene vapour pressure, and those changes can affect the solubility of the lycopene in a supercritical fluid. The relative effects of solvent density changes and solute vapour pressure may determine the degree that lycopene solubility is enhanced or impeded. The solute vapour pressure has an impact on the solubility because a crystalline solid has a lower solubility than an amorphous solid owing to the difference in free energy, and a relatively higher enthalpy of fusion (Hansen et al., 2001). An increase in temperature raises the solvent’s vapour pressure and decreases the SCF density, but the magnitude of the change in density becomes smaller at elevated temperature (Marentis, 1998). This behaviour may be explained by a complex balancing effect between the density of the solvent and the solute vapour pressure as the temperature is increased (Spanos et al., 1993). At a constant density, solubility increases with a rise in temperature, owing to the increase in vapour pressure of the solid (Johannsen and Brunner, 1997). Rozzi et al. (2002) and Hansen et al. (2001) suggest that the uncertainty of solubility measurements increases as the temperature decreases. The solubility of lycopene increased as the temperature rose at a constant pressure. The degree of increase in solubility of lycopene extracts varied over different temperature ranges. For example, the increase in solubility when the operating temperature rose from 45 to 55 °C, was less than the solubility increase from 55 to 65 °C. Moreover, the increase in solubility of lycopene extracts in CO2 fluid when the temperature increased from 55 to 65 °C was higher than the increase in solubility when from 65 to 75 °C (Shi et al., 2009a). As shown in Fig. 20.6, the solubility of lycopene from 45 to 55 °C increased along with pressure. A crossover point was noted and the magnitude of solubility increase at 30 MPa with temperature was lower than that at 20 and 25 MPa. Topal et al. (2006) extracted lycopene from dried tomato skins using supercritical CO2 and observed a similar trend. Co-solvents and modifiers Extraction can be enhanced using co-solvent or modifier that is able to interact with the target compounds. The addition of an appropriate cosolvent or modifier also influences the solubility of target compounds in
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638 Separation, extraction and concentration processes the supercritical fluid. Although the nature of CO2 drastically reduces or eliminates the solubility of many polar molecules, this is easily remedied by the addition of polar co-solvents or modifiers. Modifiers, such as methanol, increase the solubility of polar compounds but may also increase the critical temperature of the solvent. This may be a problem for compounds that are thermally labile, thus careful selection of supercritical CO2 incorporating an appropriate co-solvent or modifier is required. Another possible drawback of addition of modifiers is product contamination. This may be solved by the addition of a separation step subsequent to the extraction process (Hansen et al., 2001). Chandra and Nair (1997) indicated that the addition of co-solvents or modifiers to supercritical CO2 increase the solvating power. Chang and Randolph (1989) observed that the solubility of b-carotene increased with the addition of toluene. In their study, the results showed that the highest solubility was more than nine times higher than the extraction without toluene under the same conditions, thus showing a large b-carotene solubility enhancement. The influence of various co-solvents and modifiers at a constant temperature, pressure, and solvent flow rate was studied by Ollanketo et al. (2001). Ethanol is the most viable modifier that may be used in largescale lycopene extraction and other food applications. Cygnarowicz et al., (1990) also reported that ethanol functioned as a more effective modifier than methanol or methylene chloride. Modifiers that are miscible in water give higher recoveries. However, a disadvantage of using ethanol as a cosolvent is that an additional step must be added to the process in order to remove it from the final product. This procedure requires the use of heat, thus posing the risk of isomerization of lycopene and other carotenoids (Sun and Temelli, 2006). Natural co-solvents or modifiers used in the extraction of carotenoids include vegetable oils. Sovová et al. (2001) indicated that the modifier effect of vegetable oil was not as great as ethanol, with the increase in b-carotene solubility being four times smaller. The effect of vegetable oil as a modifier is limited owing to its low solubility in CO2. Unlike vegetable oil, canola oil containing a high degree of unsaturated fatty acids compared with the vegetable oil, has been suggested as a natural co-solvent, or modifier, and it is much more soluble in supercritical CO2 than b-carotene is; results show that the carotenoid concentration nearly doubled with canola oil addition. In addition, the effect of increasing the pressure was heightened by this modifier. Sun and Temelli (2006) attributed this effect to interactions between the modifier and the solid material. Penetration of the oil into the plant material matrix may stretch the cell walls, thus making it easier for the supercritical CO2 to penetrate the matrix. In addition, the oil may help release carotenoids in the matrix by loosening the cell structures and increasing solute exposure.
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Supercritical-fluid extraction of lycopene from tomatoes 639
20.7 Conclusion and future trends One of the most important trends in the food industry is the demand for all-natural food ingredients that are free of chemical additives. Natural food antioxidants are derivatives of plant by-products. SFE with CO2 is a viable alternative to the conventional solvent extraction technique to extract bioactive components from agricultural materials. Supercritical CO2 is nontoxic, safe, inexpensive, and has a lower critical temperature (31 °C) that can be easily reached. Therefore, SFE leaves a safe and non-toxic residue in the extracts whereas conventional methods use organic solvents and high temperatures. The mild conditions that accompany with SFE allow the extraction of lycopene and other thermally labile bioactive components from natural sources with minimal damage to the integrity or stability of lycopene and other carotenoids. Large-scale SFE has become a practicable process for the extraction of lycopene from tomato materials. The solvating power of supercritical fluids is sensitive to temperature and pressure changes, thus the extraction parameters may be optimized to provide the highest possible recoveries of lycopene with maximum antioxidant activity. Low viscosity and high mass-transfer rates are additional advantages of using SFE for food products. Ethanol and edible oil as co-solvents or modifiers have been incorporated in lycopene extraction to enhance extraction efficiency and are viable for use in the food industry. SFE offers a unique advantage of adding value to agricultural waste by extracting the lycopene from tomato skins and using it for the fortification of foods and other applications. Its drawbacks are the difficulties in extracting polar compounds and its susceptibility to extracting compounds from a complex matrix where the phase interaction with the intrinsic properties of the product inhibits its effectiveness. Some drawbacks can be ameliorated by using co-solvents. However, much additional investigation is required to understand the solvation effects on targeted bioactive components being extracted. There are many factors which influence the solubility of lycopene and extraction yield in SFE. The fluid (e.g. CO2) density, operating pressure and temperature, as well as flow rate of fluid affect the quality of the extracts and the efficiency of the extraction. By understanding the effect of the parameters that influence extraction, conditions can be chosen to optimize the recovery and cost efficiency. When determining the parameters that should be used to maximize the recovery of lycopene and the solubility, many studies attempted to optimize the operating conditions with incorporated modifier or co-solvent that may be applicable in large-scale applications. For example, non-toxic co-solvents and modifiers that would be acceptable for food processing. Therefore, canola oil and ethanol was used to enhance the extraction yields. The nature of the material used as a source of lycopene governs the availability of the compounds for the extraction process. Although high temperatures in the extraction process generally increase the solubilities of components in supercritical fluids, the conditions under
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640 Separation, extraction and concentration processes which thermally labile targeted compounds are negatively affected should be considered. The intensity and the length of heat processing affect the degree of lycopene isomerization, thus ideally the extraction time and temperature should be minimized. Minimizing such conditions also leads to a more economically viable process. Excessively high flow rates may reduce the contact time between the target components and the solvent and restrict the fluid flow in the samples if it becomes compacted. The optimal flow rate appears to vary with the targeted molecule, the relatively high flow rates having a negative effect on some components. Therefore, to successfully implement SFE technology on an industrial scale, it is necessary to understand the technology, focusing on the proper design of plant components and optimization of the process parameters that provide maximum recovery and quality with minimize operating costs. SFE technology is available in the form of a single-stage batch process, which could be upgraded to a multistage semi-continuous batch, coupled with a multi-separation process to reduce extraction time and increase extraction efficiency. The need to improve the design into continuous modes is growing. Although the capital cost of SFE plant is higher than the traditional or conventional organic solvent extraction plant, the operating costs are lower. However, for a proper comparison of the capital cost of SFE plant versus conventional extraction plant, it is necessary to take into account all the associated equipment used in the conventional extraction processes, such as distillation or evaporation for removal of solvent from the final extracts. On the other hand, the selling price of the products obtained by SFE are also higher because the extracts are safe and natural without chemical contamination. Further development of SFE is necessary for largescale production to be cost effective. Therefore, with improved processing conditions and reduced cost, lycopene extraction from tomato materials by SFE should become even more economical at low throughputs.
20.8 References Anitescu, G., Doneanu, C. and Radulescu, V. 1997. Isolation of coriander oil: comparison between steam distillation and supercritical CO2 extraction, Flavour and Fragrance Journal, 12, 173–176. Arab, L. and Steck, S. 2000. Lycopene and cardiovascular disease. American Journal of Clinical Nutrition, 71, 1691–1695. Baysal, T., Ersus, S. and Starmans, J. D. A. 2000. Supercritical CO2 extraction of b-carotene and lycopene from tomato paste waste. Journal of Agriculture and Food Chemistry, 48, 5507–5511. Boileau, A. C., Merchen, N. R., Wasson, K., Atkinson, C. A. and Erdman, J. W. 1999. Cis-lycopene is more bioavailable than trans-lycopene in vitro and in vivo in lymphcannulated ferrets. Journal of Nutrition, 129, 1176–1181. Boon, C. S., Xu, Z., Yue, X., McClements, D. J., Weiss, J. and Decker, E. A. 2008. Factors affecting lycopene oxidation in oil-in-water emulsions. Journal of Agricultural and Food Chemistry, 56(4), 1408–1414.
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Supercritical-fluid extraction of lycopene from tomatoes 641 Boskovic, M. A. 1979. Fate of lycopene in dehydrated tomato products: carotenoid isomerization in food system. Journal of Food Science, 44, 84–86. Brady, B. O., Chien-Ping, C. K., Dooley, K. M., Knopf, F. C. and Gambrell, R. P. 1987. Supercritical extraction of toxic organics from soils. Industrial & Engineering Chemistry Research, 26, 261–268. Breinholt, B., Lauridsen, S. T., Daneshvar, B. and Jakobsen, J. 2000. Dose–response effects of lycopene on selected drug-metabolizing and antioxidant enzymes in the rat, Cancer Letters, 154, 201–210. Cadoni, E., Giorgi, M. R., Medda, E. and Poma, G. 2000. Supercritical CO2 extraction of lycopene and b-carotene from ripe tomatoes. Dyes and Pigments, 44, 27–32. Calvey, E. M. and Page, S. W. 1990. Apparent solubility threshold densities of substituted coumarins, Journal of Supercritical Fluids, 3, 115–120. Casas, L., Mantell, C., Rodríguez, M., Torres, A., Macías, F.A. and Martínez de la Ossa, E. 2007. Effect of the addition of cosolvent on the supercritical fluid extraction of bioactive compounds from Helianthus annuus L. Journal of Supercritical Fluids, 41, 43–49. Chandra, A. and Nair, M. G. 1997. Supercritical fluid carbon dioxide extraction of a- and b-carotene from carrot (Daucus carota L.). Phytochemical Analysis, 8, 244–246. Chang, C. J. and Randolph, A. D. 1989. Precipitation of microsize organic particles from supercritical fluids. AIChE Journal, 35(11), 1879–1882. Chao, R. R., Mulvaney, S. J. and Hanah, H. 1993. Effects of extraction and fractionation pressures on supercritical extraction of cholesterol from beef tallow, Journal of the American Oil Chemists’ Society, 70, 139–143. Ciurlia, L., Bleve, M. and Rescio, L. 2009. Supercritical carbon dioxide co-extraction of tomatoes (Lycopersicum esculentum L.) and hazelnuts (Corylus avellana L.): a new procedure in obtaining a source of natural lycopene. Journal of Supercritical Fluids, 49(3), 338–344 Cocero, M. J., González, S., Pérez, S. and Alonso, E. 2003. Supercritical extraction of unsaturated products. Degradation of b-carotene in supercritical extraction process, Journal of Supercritical Fluids, 19, 39–44. Clinton, S. K. 1998. Lycopene: chemistry, biology, and implications for human health and disease. Nutrition Reviews, 56(2), 35–51. Cygnarowicz, M. L., Maxwell, R. J. and Seider, W. D. 1990. Equilibrium solubilites of b-carotene in supercritical carbon dioxide. Fluid Phase Equilibria, 59, 57–71. Dewanto, V., Wu, X., Adom, K. K. and Liu, R. H. 2002. Thermal processing enhances the nutritional value of tomatoes by increasing total antioxidant activity. Journal of Agriculture and Food Chemistry, 50, 3010–3014. Ferreira, S. R. S. and Meireles, M. A. A. 2002. Modeling the supercritical fluid extraction of black pepper (Piper nigrum L.) essential oil. Journal of Food Engineering, 54, 263–269. Franca, L. F. and Meireles, M. A. A. 2000. Modeling the extraction of carotene and lipids from pressed palm oil (Elaes guineensis) fiber using supercritical CO2. Journal of Supercritical Fluids, 18, 35–47. Friedrich, J. P. and Pryde, E. H. 1984. Supercritical CO2 extraction of lipid-bearing materials and characterization of products, Journal of the American Oil Chemists Society, 61, 223. Gordon, H. T. and Bauernfeind, J. C. 1982. Carotenoids as food colorants. Critical Reviews in Food Science and Nutrition, 18, 59–97. Gouveia, L., Nobre, B. P., Marcelo, F. M., Mrejen, S., Cardoso, M. T., Palavra, A. F. and Mendes, R. L. 2006. Functional food oil coloured by pigments extracted from microalgae with supercritical CO2. Food Chemistry, 43, 2876–2878. Güçlü-Üstündag, Ö. and Temeli, F. 2005. Solubility behavior of ternary systems of lipids, cosolvents and supercritical carbon dioxide and processing aspects. Journal of Supercritical Fluids, 36, 1–15.
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642 Separation, extraction and concentration processes Hackett, M. M., Lee, J. H., Francis, D. and Schwartz, S. J. 2004. Thermal stability and isomerisation of lycopene in tomato oleoresins from different varieties. Journal of Food Science, 69, 536–541. Hadley, C. W., Clinton, S. K. and Schwartz, S. J. 2003. The consumption of processed tomato products enhances plasma lycopene concentrations in association with reduced lipoprotein sensitivity to oxidative damage. Journal of Nutrition, 133, 727–732. Hansen, B. N., Harvey, A. H., Coelho, J. A. P., Palavra, A. M. F. and Bruno, T. J. 2001. Solubility of capsaicin and b-carotene in supercritical carbon dioxide and in halocarbons. Journal of Chemical and Engineering Data, 46, 1054–1058. Hawthorne, S. B., Galy, A. B., Schmitt, V. O. and Miller, D. J. 1995. Effect of SFE flow rate on extraction rates: classifying sample extraction behavior. Analytical Chemistry, 67, 2723–2732. Hopper, M. L. and King, J. W. 1991. Enhanced supercritical fluid carbon dioxide extraction of pesticides from foods using palletized diatomaceous earth, Journal of the Association of Official Analytical Chemists, 7, 661–666. Johannsen, M. and Brunner, G. 1997. Solubilities of the fat-soluble vitamins A, D, E, and K in supercritical carbon dioxide. Journal of Chemical and Engineering Data, 42, 106–111. Joslin, C. G., Gray, C. G. and Goldman, S. 1996. Solubility in supercritical fluids from the virial equation of state, Molecular Physics, 89, 489–503. Kassama, L. S., Shi, J. and Mittal, G. S. 2008. Optimization of supercritical fluid extraction of lycopene from tomato skin with central composite rotatable design model, Separation and Purification Technology, 60, 278–284. Krichnavaruk, S., Shotipruk, A., Goto, M. and Pavasant, P. 2008. Supercritical carbon dioxide extraction of astaxanthin from Haematococcus pluvialis with vegetable oils as co-solvent. Bioresource Technology, 99, 5556–5560. Krinsky, N. I. and Rock, C. L. 1998. Carotenoids: chemistry, sources and physiology. In M. Sadler, S. Strain, B. Caballero (Eds.), Encyclopedia of Human Nutrition (pp. 304–314). Academic Press, London, UK. Lehotay, S. J. 1997. Supercritical fluid extraction of pesticides in foods. Journal of Chromatography A, 785, 289–312. Liu, L. H., Zabaras D., Bennett, L. E., Aguas, P. and Woonton, P. W. 2009. Effects of UV-C, red light and sun light on the carotenoid content and physical qualities of tomatoes during post-harvest storage. Food Chemistry, 115, 495–500. Luque de Castro, M. D., Valcarcel, M. and Tena, M. T. 1994. Analytical supercritical fluid extraction. Springer-Verlag, Berlin, Germany. Marentis, R. T. 1988. Steps to developing a commercial supercritical carbon dioxide processing plant. In Supercritical fluid extraction and chromatography. Charpentier, B. A. and Sevenants, M. R., Ed.; American Chemical Society, Symposium Series, 127–143. Marsili, R. and Callahan, D. 1993. Comparison of a liquid solvent extraction technique and supercritical fluid extraction for the determination of a- and b-carotene in vegetable. Journal of Chromatographic Science, 31, 422–428. Mascio, P. D., Kaiser, S. and Sies, H. 1989. Lycopene as the most efficient biological carotenoid singlet oxygen quencher. Archives of Biochemistry and Biophysics, 274, 532–538. Mayeaux, M., Xu, Z., King, J. M. and Prinyawiwatkul, W. 2006. Effects of cooking conditions on the lycopene content in tomatoes. Journal of Food Science, 71, 461–464. Mayer-Miebach, E., Behsnilian, D., Regier, M. and Schuchmann, H. P. 2005. Thermal processing of carrots: lycopene stability and isomerisation with regard to antioxidant potential. Food Research International, 38, 1103–1108. Mayer-Miebach, E. and Spiess, W. E. L. 2003. Influence of cold storage and blanching on the carotenoid content of Kintoki carrots. Journal of Food Engineering, 65, 211–213. © Woodhead Publishing Limited, 2010
Supercritical-fluid extraction of lycopene from tomatoes 643 Nagy, B. and Simándi, B. 2008. Effects of particle size distribution, moisture content, and initial oil content on the supercritical fluid extraction of paprika. Journal of Supercritical Fluids, 46(3), 293–298. Negre-Salvayre, A., Dousset, N., Ferretti, G., Bacchetti, T., Curatola, G. and Salvayre, R. 2006. Antioxidant and cytoprotective properties of high-density lipoproteins in vascular cells. Free Radical Biology and Medicine, 41, 1031–1040. O’Day, D. M. and Rosenau, J. R. 1982. Solvent extraction of carotenoids from alfalfa. Transactions of the ASAE (American Society of Agricultural Engineers), 25, 515– 519. Ollanketo, M., Hartonen, K., Riekkola, M. L., Holm, Y. and Hiltunen, R. 2001. Supercritical carbon dioxide extraction of lycopene in tomato skins. European Food Research and Technology, 212, 561–565. Olson, J. 1986. Carotenoid, vitamin A and cancer. Journal of Nutrition, 116, 1127– 1130. Palozza, P. 1998. Prooxidant actions of carotenoids in biologic systems. Nutrition Reviews, 56(9), 257–265. Papas, A. 1999. Diet and antioxidants status: In Antioxidant status, diet, nutrition and health. Ed. Andreas M. Papas. pp. 89–106. CRC Press, New York. Pol, J., Hyotylainen, T., Ranta-Aho, O. and Riekkola, M-L. 2004. Determination of lycopene in food by on-line SFE coupled to HPLC using a single monolithic column for trapping and separation. Journal of Chromatography A, 1052, 25–31. Pourmortazavi, S. M. and Hajimirsadeghi, S. S. 2007. Supercritical fluid extraction in plant essential and volatile oil analysis. Journal of Chromatograph A, 1163, 2–24. Rao, A. V. and Agarwal, S. 1998. Bioavailability and in vivo antioxidant properties of lycopene from tomato products and their possible role in the prevention of cancer. Nutrition and Cancer. 31, 199–203. Rao, A. V. and Agarwal, S. 1999. Role of lycopene as antioxidant carotenoid in the prevention of chronic diseases: a review. Nutrition Research, 19(2), 305–32. Rao, A.V. and Rao, L.G. 2007. Carotenoids and human health. Pharmacological Research, 55(3), 207–216. Rozzi, N. L., Singh, R. K., Vierling, R. A. and Watkins, B. A. 2002. Supercritical fluid extraction of lycopene from tomato processing by-products. Journal of Agriculture and Food Chemistry, 50, 2638–2643. Sabio, E., Lozano, M., Montero de Espinosa, V., Mendes, R. L., Pereira, A. P., Palavra, A. F. and Coelho, J. A. 2003. Lycopene and b-carotene extraction from tomato processing waste using supercritical CO2. Industrial & Engineering Chemistry Research, 42, 6641–6646. Sakaki, K. 1992. Solubility of b-carotene in dense carbon dioxide and nitrous oxide from 308 to 323 K and from 9.6 to 30 MPa. Journal of Chemical and Engineering Data, 37, 249–251. Seybold, C., Frohlich, K., Bitsch, R., Otto, K. and Bohm, V. 2004. Changes in contents of carotenoids and vitamin E during tomato processing. Journal of Agriculture and Food Chemistry, 52, 7005–7010. Shi, J. 2002. Lycopene: biochemistry and functionality. Food Science and Biotechnology, 11(5), 574–581. Shi, J., Bryan, M., Le Maguer, M. and Kakuda, Y. 2002a. Kinetics of lycopene degradation in tomato puree with heat and light irradiation treatment. Journal of Food Processing Engineering, 25, 485–498. Shi, J., Dai, Y., Kakuda, Y., Mittal, G. and Xue, J. 2008. Effect of heating and light irradiation on the stability of lycopene in tomato purée. Food Control, 19(5), 514–520. Shi, J., Kakuda, Y., Zhou, X., Mittal, G. and Pan, Q. 2007a. Correlation of mass transfer coefficient in the extraction of plant oil in a fixed bed for supercritical CO2. Journal of Food Engineering, 78, 33–40. Shi, J., Kakuda Y. and Yeung, D. 2004a. Antioxidative properties of lycopene and other carotenoids: synergistic effects. BioFactors, 21, 203–210. © Woodhead Publishing Limited, 2010
644 Separation, extraction and concentration processes Shi, J., Kassama, L. and Kakuda, Y. 2006. Supercritical fluid technology for extraction of bioactive components. Functional food ingredients and nutraceuticals: processing technology, Ed. J. Shi. CRC Press, USA. p. 45–73. Shi, J., Khatri, M., Xue, J., Mittal, G., Ma, Y. and Li, D. 2009a. Solubility of lycopene in supercritical CO2 fluid affected by temperature and pressure. Separation and Purification Technology, 66, 322–328. Shi, J., Le Maguer, M. and Kakuda, Y. 1999a. Lycopene degradation and isomerization in tomato dehydration. Food Research International, 32, 15–21. Shi, J., Le Maguer, M. and Wang, S. 1999b. Chemical composition of tomatoes affected by maturity and fertility practices. Journal of Food Quality, 22, 147–156. Shi, J., MacNaughton, L., Kakuda, Y., Bettger, W., Yeung, D. and Jiang, Y. 2004c. Bioavailability of lycopene from tomato products. Journal of Food Science and Nutrition, 9, 98–106. Shi, J., Mittal, G., Kim, E. and Xue, S. 2007b. Solubility of carotenoids in supercritical CO2. Food Review International, 23, 341–371. Shi, J., Qu, Q., Kakuda, Y., Xue, J., Jiang, Y., Koide, S. and Shim, Y. 2007c. Investigation of the antioxidant and synergistic effect of lycopene and vitamin E on the AMVN induced oxidation of LAME. Journal of Food Composition and Analysis, 20(6), 603–608. Shi, J., Qu, Q., Kakuda, Y. and Yeung, D. 2004c. Stability and bioavailability and synergistic effect of antioxidant capacity of lycopene with other antioxidants. Critical Reviews in Food Science and Nutrition, 44, 559–573. Shi, J., Wu, Y., Bryan, M. and Le Maguer, M. 2002b. Oxidation and isomerization of lycopene under thermal treatment and light irradiation in food processing. Nutraceuticals and Foods, 7(2), 179–183. Shi, J., Yi, C., Xue, J., Jiang, Y., Ma, Y. and Li, D. 2009b. Effects of modifier on lycopene extract profile from tomato skin using supercritical-CO2 fluid. Journal of Food Engineering, 93, 431–436. Shi, J., Zhou, X. and Kassama, L. 2007d. Correlation of mass transfer coefficient in separation process with supercritical CO2. Drying Technology International, 25, 335–339. Shi, J. and Le Maguer, M. 2000. Lycopene in tomatoes: chemical and physical properties affected by food processing. Critical Reviews in Food Science and Nutrition CRC, 40(1), 1–41. Shi, J. and Zhou, X. 2006. Solubility property of bioactive components on recovery yield in separation process by supercritical fluid. Functional food ingredients and nutraceuticals: processing technology, Ed. J. Shi. CRC Press, USA. p. 3–43. Shixian, Q., Dai, Y., Kakuda, Y., Shi, J., Mittal, G., Yeung, D. and Jiang,Y. 2005. Synergistic antioxidative effects of lycopene with other bioactive compounds. Food Review International, 21, 295–311. Simandi, B., Kristo, S. T., Kery, A., Selmeczi, L. K., Kmecz, I. and Kemeny, S. 2002. Supercritical fluid extraction of dandelion leaves, Journal of Supercritical Fluids, 23, 135–142. Sovová, H., Stateva, R. P. and Galushko, A. A. 2001. Solubility of b-carotene in supercritical CO2 and the effect of entrainers. Journal of Supercritical Fluids, 21, 195–203. Spanos, G. A., Chen, H. and Schwartz, S. J. 1993. Supercritical CO2 extraction of b-carotene from sweet potatoes. Journal of Food Science, 58(4), 817–820. Stahl, W. and Sies, H. 1992. Uptake of lycopene and its geometrical isomers is greater from heat-processed than from unprocessed tomato juice in humans. Journal of Nutrition, 122, 2161–2166. Stahl, W. and Sies, H. 1996. Perspectives in biochemistry and biophysics. Archives of Biochemistry and Biophysics, 336 (1), 1–9. Subra, P., Castellani, S., Ksibi, H. and Garrabos, Y. 1997. Contribution to the determination of the solubility of b-carotene in supercritical carbon dioxide and nitrous oxide, experimental data and modelling. Fluid Phase Equilibria, 131, 269–286. © Woodhead Publishing Limited, 2010
Supercritical-fluid extraction of lycopene from tomatoes 645 Suganuma, H. and Inakuma, T. 1999. Protective effect of dietary tomato against endothelial dysfunction in hypercholesterolemic mice. Bioscience, Biotechnology, and Biochemistry, 63, 78–82. Sun, M. and Temelli, F. 2006. Supercritical carbon dioxide extractions of carotenoids from carrot using canola oil as a continuous co-solvent. Journal of Supercritical Fluids, 37, 397–408. Tonthubthimthong, P., Chuapraset, S., Douglas, P. and Luewisutthicat, W. 2001. Supercritical CO2 extraction of nimbin from neem seeds – an experimental study, Journal of Food Engineering, 47, 289–293. Toor, R. K. and Savage, G. P. 2006. Effect of semi-drying on the antioxidant components of tomatoes. Food Chemistry, 94, 90–97. Topal, U., Sasaki, M., Goto, M. and Hayakawa, K. 2006. Extraction of lycopene from tomato skin with supercritical carbon dioxide: Effect of operating conditions and solubility analysis. Journal of Agriculture and Food Chemistry, 54, 5604–5610. Tsai, C., Lin, H. and Lee, M. 2007. Solubility of disperse yellow 54 in supercritical carbon dioxide with or without cosolvent, Fluid Phase Equilibria, 260, 287–294. Vági, E., Simándi, B., Vásárhelyiné, K. P., Daood, H., Kéry, Á., Doleschall, F. and Nagy, B. 2007. Supercritical carbon dioxide extraction of carotenoids, tocopherols and sitosterols from industrial tomato by-products. Journal of Supercritical Fluid, 40, 218–226. Vasapollo, G., Longo, L., Rescio, L. and Ciurlia, L. 2004. Innovative supercritical CO 2 extraction of lycopene from tomato in the presence of vegetable oil as co-solvent. Journal of Supercritical Fluids, 29, 87–96. Vaughn, K. L. S., Clausen, E. C., King, J. W., Howard, L. R. and Carrier, D. J. 2008. Extraction conditions affecting supercritical fluid extraction (SFE) of lycopene from watermelon. Bioresource Technology, 99(16), 7835–7841. Vega, P. J., Balaban, M. O., Sims, C. A., Keefe, S. F. and Cornell, J. A. 1966. Supercritical carbon dioxide extraction efficiency for carotenes from carrots by RSM, Journal of Food Science, 61, 757–765. Walsh, J. M., Ikonomou, G. D. and Donohue, M. D. 1987. Supercritical phase behavior: the entrainer effect. Fluid Phase Equilibria, 33, 295–314. Willcox, J. K., Catignani, G. and Lazarus, S. 2003. Tomatoes and cardiovascular health. Critical Reviews in Food Science and Nutrition, 43, 1–18. Wu, D. and Meydini, S. N. 1999. Antioxidants and immune function: In Antioxidant status, diet, nutrition and health. ed. Andreas M. Papas. pp. 371–400. CRC Press, New York. Xianquan, S., Shi, J., Kakuda, Y. and Jiang, Y. 2005. Stability of lycopene during food processing and storage. Journal of Medicinal Food, 8(4), 413–422. Yi, C., Shi, J., Xue, J., Jiang, Y. and Li, D. 2009. Effects of supercritical fluid extraction parameters on lycopene yield and antioxidant activity. Food Chemistry, 113(4), 1088–1094. Ziémons, E., Goffin, E., Lejeune, R., Proença da Cunha, A., Angenot, L. and Thunus, L. 2005. Supercritical carbon dioxide extraction of tagitinin C from Tithonia diversifolia. Journal of Supercritical Fluids, 33, 53–59.
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Lycopene concentration (mg g–1)
32.00 26.25 20.50 14.75 9.00 40
100 35 70
30
Lycopene concentration (mg g–1)
Pressure (MPa)
25
20 40 (a)
55
85
Temperature (°C)
32.00 26.25 20.50 14.75 9.00 2.00 1.75
85
1.50
Flow rate (mL min–1)
1.25
40
55
100
70 Temperature (°C)
Lycopene concentration (mg g–1)
(b)
32.00 26.25 20.50 14.75 9.00 2.00 1.75 1.50
Flow rate (mL min–1)
1.25
1.00 20 (c)
Plate II Three-dimensional response surface plots showing lycopene yield from SFE: (a) effects of temperature and pressure at constant flow rate (1.5 mL min–1); (b) effects of temperature and flow rate 40 at constant pressure (30 35 MPa); (c) effects of pressure 30 and flow rate at constant 25 Pressure (MPa) temperature (70 °C) (modified from Yi et al., 2009).
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Separation, extraction and concentration processes in the food, beverage and nutraceutical industries
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Related titles: Separation processes in the food and biotechnology industries (ISBN 978-1-85573-287-2) This book reviews methods and techniques for separating food components and products of the biotechnology industry. The introduction focuses on food composition and some of the conventional separation techniques. Subsequent chapters deal with each specific type or area of application individually and include information on the basic principles, industrial equipment available, commercial applications and an overview of research and development. Novel enzyme technology for food applications (ISBN 978-1-84569-132-5) The food industry is constantly seeking advanced technologies to produce valueadded, nutritionally-balanced products for consumers in a sustainable fashion. Since enzymes are so specific in their action, they are a useful biotechnological processing tool and by controlling the action of enzymes, innovative food ingredients and higher quality food products can be produced. Part one of Novel enzyme technology for food applications covers the principles of industrial enzyme technology, including methods to develop and tailor enzymes for food bioprocessing. Part two introduces the reader to novel applications of enzymes for the production of improved ingredients and food products. Food processing technology (Third edition) (ISBN 978-1-84569-216-2) The first edition of Food processing technology was quickly adopted as the standard text by many food science and technology courses. The publication of a completely revised and updated third edition consolidates the position of this textbook as the best single-volume introduction to food manufacturing technologies available. The third edition has been updated and extended to include the many developments that have taken place since the second edition was published. In particular, advances in microprocessor control of equipment, ‘minimal’ processing technologies, functional foods, developments in ‘active’ or ‘intelligent’ packaging, and storage and distribution logistics are described. Technologies that relate to cost savings, environmental improvement or enhanced product quality are highlighted. Additionally, sections in each chapter on the impact of processing on food-borne micro-organisms are included for the first time. Details of these and other Woodhead Publishing books can be obtained by: ∑ visiting our web site at www.woodheadpublishing.com ∑ contacting Customer Services (e-mail: [email protected]; fax: +44 (0) 1223 893694; tel.: +44 (0) 1223 891358 ext. 130; address: Woodhead Publishing Limited, Abington Hall, Granta Park, Great Abington, Cambridge CB21 6AH, UK) If you would like to receive information on forthcoming titles, please send your address details to: Francis Dodds (address, tel. and fax as above; e-mail: francis. [email protected]). Please confirm which subject areas you are interested in.
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iii Woodhead Publishing Series in Food Science, Technology and Nutrition: Number 202
Separation, extraction and concentration processes in the food, beverage and nutraceutical industries Edited by Syed S. H. Rizvi
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iv Published by Woodhead Publishing Limited, Abington Hall, Granta Park, Great Abington, Cambridge CB21 6AH, UK www.woodheadpublishing.com Woodhead Publishing, 525 South 4th Street #241, Philadelphia, PA 19147, USA Woodhead Publishing India Private Limited, G-2, Vardaan House, 7/28 Ansari Road, Daryaganj, New Delhi – 110002, India www.woodheadpublishingindia.com First published 2010, Woodhead Publishing Limited © Woodhead Publishing Limited, 2010 The authors have asserted their moral rights. This book contains information obtained from authentic and highly regarded sources. Reprinted material is quoted with permission, and sources are indicated. Reasonable efforts have been made to publish reliable data and information, but the authors and the publisher cannot assume responsibility for the validity of all materials. Neither the authors nor the publisher, nor anyone else associated with this publication, shall be liable for any loss, damage or liability directly or indirectly caused or alleged to be caused by this book. Neither this book nor any part may be reproduced or transmitted in any form or by any means, electronic or mechanical, including photocopying, microfilming and recording, or by any information storage or retrieval system, without permission in writing from Woodhead Publishing Limited. The consent of Woodhead Publishing Limited does not extend to copying for general distribution, for promotion, for creating new works, or for resale. Specific permission must be obtained in writing from Woodhead Publishing Limited for such copying. Trademark notice: Product or corporate names may be trademarks or registered trademarks, and are used only for identification and explanation, without intent to infringe. British Library Cataloguing in Publication Data A catalogue record for this book is available from the British Library. ISBN 978-1-84569-645-0 (print) ISBN 978-0-85709-075-1 (online) ISSN 2042-8049 Woodhead Publishing Series in Food Science, Technology and Nutrition (print) ISSN 2042-8057 Woodhead Publishing Series in Food Science, Technology and Nutrition (online)
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Contents
Contributor contact details..................................................................
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Woodhead Publishing Series in Food Science, Technology and Nutrition............................................................................................... xvii Preface................................................................................................. xxvii Part I Developments in food and nutraceutical separation, extraction and concentration techniques 1 Principles of supercritical fluid extraction and applications in the food, beverage and nutraceutical industries................. Ž. Knez, M. Škerget and M. Knez Hrnčič, University of Maribor, Slovenia 1.1 Introduction....................................................................... 1.2 Thermodynamic fundamentals.......................................... 1.3 Cycle processes for extraction using supercritical fluids 1.4 Extraction of solids using SCF......................................... 1.5 Extraction of liquids using SCF........................................ 1.6 Conclusion......................................................................... 1.7 References......................................................................... 2 Principles of pressurized fluid extraction and environmental, food and agricultural applications................. C. Turner and M. Waldebäck, Uppsala University, Sweden 2.1 Introduction....................................................................... 2.2 Instrumentation and principles of pressurized fluid extraction . ........................................................................ 2.3 Applications of pressurized fluid extraction..................... 2.4 Future trends..................................................................... 2.5 Sources of further information and advice....................... 2.6 Conclusions ...................................................................... 2.7 References......................................................................... © Woodhead Publishing Limited, 2010
3 3 8 21 26 30 32 36 39 39 41 56 59 61 63 64
vi Contents 3 Principles of physically assisted extractions and applications in the food, beverage and nutraceutical industries..................................................................................... 71 E. Vorobiev, Compiègne University of Technology, France and F. Chemat, University of Avignon and Pays de Vaucluse, France 3.1 Introduction....................................................................... 71 3.2 Pulsed electric field-assisted extractions in the food industry.............................................................................. 72 3.3 Ohmic heating-assisted extractions in the food industry 83 3.4 Extraction assisted by high-voltage electrical discharges and applications in the food industry................................ 86 3.5 Ultrasound-assisted extraction (UAE) in the food industry.............................................................................. 90 3.6 Microwave-assisted extraction (MAE) in the food industry.............................................................................. 96 3.7 Combination of physical treatments for extraction in the food industry............................................................... 100 3.8 References......................................................................... 102 4 Advances in process chromatography and applications in the food, beverage and nutraceutical industries .................... M. Ottens and S. Chilamkurthi, Delft University of Technology, The Netherlands 4.1 Introduction....................................................................... 4.2 Basic principles of process chromatography.................... 4.3 Applications of process chromatography in the food, beverage and nutraceutical industries............................... 4.4 Recent developments in process chromatography............ 4.5 Process control in chromatography................................... 4.6 Future trends..................................................................... 4.7 Conclusions....................................................................... 4.8 Sources of further information and advice....................... 4.9 List of abbreviations......................................................... 4.10 References......................................................................... 5 Novel adsorbents and approaches for nutraceutical separation.................................................................................... B. W. Woonton, CSIRO Food and Nutritional Sciences, Australia and G. W. Smithers, Food Industry Consultant, Australia 5.1 Introduction....................................................................... 5.2 Molecular imprinted polymers and applications in the nutraceutical industry........................................................ 5.3 Organic monoliths and applications in the nutraceutical industry..............................................................................
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148 149 153
Contents vii 5.4 5.5 5.6 5.7 5.8 5.9
Stimuli-responsive resins and applications in the nutraceutical industry........................................................ Mesoporous molecular sieves and applications in the nutraceutical industry........................................................ Peptide affinity ligands and phage display methodology and applications in the nutraceutical industry................... Membrane adsorbers, membrane chromatography and applications in the nutraceutical industry......................... Conclusions and sources of further information and advice................................................................................ References.........................................................................
6 Advances in the effective application of membrane technologies in the food industry.............................................. M. Pinelo, G. Jonsson and A. S. Meyer, Technical University of Denmark, Denmark 6.1 Introduction....................................................................... 6.2 Theoretical fundamentals of membrane separation.......... 6.3 Membrane technology in the dairy industry..................... 6.4 Membrane technology in the fruit juice industry............. 6.5 Membrane technology for treatment of wastewater in the food industry............................................................... 6.6 New applications of membrane technology for the food industry: concentration and fractionation of saccharides 6.7 Future trends..................................................................... 6.8 References......................................................................... 7 Electrodialytic phenomena, associated electromembrane technologies and applications in the food, beverage and nutraceutical industries............................................................. L. Bazinet, A. Doyen and C. Roblet, Laval University, Canada 7.1 Introduction....................................................................... 7.2 Principles of electrodialytic phenomena and associated membrane technologies..................................................... 7.3 Applications of electrodialytic phenomena and associated membrane technologies................................... 7.4 Future trends .................................................................... 7.5 References......................................................................... 8 Principles of pervaporation for the recovery of aroma compounds and applications in the food and beverage industries..................................................................................... S. Sahin, Middle East Technical University, Turkey 8.1 Introduction....................................................................... 8.2 Principles of pervaporation............................................... 8.3 Transport mechanism in pervaporation for the recovery of aroma compounds......................................................... © Woodhead Publishing Limited, 2010
159 163 166 169 172 173 180 180 181 182 185 190 191 195 197
202 202 203 204 213 214
219 219 220 221
viii Contents 8.4 8.5 8.6 8.7
Selection of membranes for pervaporation in the recovery of aroma compounds.......................................... Recovery of aroma compounds by pervaporation and applications in the food and beverage industries.............. Sources of further information and future trends............. References.........................................................................
9 Advances in membrane-based concentration in the food and beverage industries: direct osmosis and membrane contactors ................................................................................... E. Drioli and A. Cassano, Institute on Membrane Technology, ITM-CNR, Italy 9.1 Introduction....................................................................... 9.2 Conventional technologies in the food and beverage industries........................................................................... 9.3 Direct osmosis and applications in the food and beverage industries............................................................ 9.4 Membrane contactors and applications in the food and beverage industries............................................................ 9.5 Conclusions....................................................................... 9.6 Nomenclature.................................................................... 9.7 References......................................................................... 10 Separation of value-added bioproducts by colloidal gas aphrons (CGA) flotation and applications in the recovery of value-added food products.................................................... P. Jauregi and M. Dermiki, The University of Reading, UK 10.1 Introduction ...................................................................... 10.2 Colloidal gas aphrons (CGA) properties........................... 10.3 Applications of CGA in the recovery of value-added food products..................................................................... 10.4 Feasibility of industrial application of CGA..................... 10.5 Future trends..................................................................... 10.6 Sources of further information and advice....................... 10.7 References......................................................................... 11 Membrane bioreactors and the production of food ingredients................................................................................... M.-P. Belleville, D. Paolucci-Jeanjean and G. M. Rios, European Institute of Membranes, France 11.1 Introduction....................................................................... 11.2 Membrane bioreactors for the production of food ingredients......................................................................... 11.3 Applications of membrane bioreactors in food industries. 11.4 Future trends..................................................................... 11.5 References.........................................................................
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227 230 239 240
244 244 245 248 250 275 275 278
284 284 285 293 307 308 309 310 314 314 315 322 331 331
Contents ix Part II Separation technologies in the processing of particular foods and nutraceuticals 12
Separation technologies in dairy and egg processing............. G. Gésan-Guiziou, INRA, France 12.1 Introduction....................................................................... 12.2 The dairy industry and composition of dairy products..... 12.3 Pretreatment of milk using separation techniques............ 12.4 Standardization and concentration of milk proteins in the dairy industry.............................................................. 12.5 Isolation of whole casein in the dairy industry................. 12.6 Separation techniques applied to whey and derivatives in the production of cheese............................................... 12.7 Fractionation of individual proteins and peptides in the dairy industry.................................................................... 12.8 Treatment of effluents and technical fluids in the dairy industry.............................................................................. 12.9 Conclusions and future trends in the dairy industry......... 12.10 The egg products industry and composition of egg products............................................................................. 12.11 Concentration and stabilization of egg white and whole egg..................................................................................... 12.12 Industrial extraction of egg-white proteins....................... 12.13 Industrial extraction of yolk components......................... 12.14 Conclusions and future trends in the egg-processing industry.............................................................................. 12.15 Sources of further information and advice....................... 12.16 References.........................................................................
13
341 341 343 347 351 354 357 360 366 368 369 371 371 374 375 376 377
Separation technologies in the processing of fruit juices....... G. Vatai, Corvinus University of Budapest, Hungary 13.1 Introduction....................................................................... 13.2 Characteristics of foods/fluids in the fruit juice product sector................................................................................. 13.3 Designing separation processes to optimize product quality in the fruit juice product sector............................. 13.4 Production of fruit juice concentrate ............................... 13.5 References.........................................................................
381
1 4
396
Separation technologies in oilseed processing ........................ M. A. Williams, Anderson International Corp., USA 14.1 Introduction....................................................................... 14.2 Preparation for oilseed processing.................................... 14.3 Extrusion preparation for oilseed processing.................... 14.4 Mechanical pressing of oilseeds....................................... 14.5 Percolation solvent extraction in oilseed processing........ 14.6 Solvent recovery in oilseed processing............................. © Woodhead Publishing Limited, 2010
381 382 383 386 394
396 397 399 403 415 422
x Contents
14.7 14.8 14.9 14.10
Obtaining oil from fruit pulps........................................... Future trends..................................................................... Sources of further information and advice....................... References.........................................................................
424 425 427 428
1 5
Separation technologies in brewing.......................................... G. J. Freeman, Campden BRI, UK 15.1 Introduction....................................................................... 15.2 Characteristics of brewery products.................................. 15.3 Selection of technology and raw materials appropriate to brewery products........................................................... 15.4 Wort production in the brewhouse.................................... 15.5 Whirlpools and applications in brewing........................... 15.6 Yeast flocculation and applications in brewing................ 15.7 Beer fining agents ............................................................ 15.8 Filter aid filtration and applications in brewing................ 15.9 Regenerable and reusable filter aids and applications in brewing......................................................................... 15.10 Bulk beer filtration by membranes.................................... 15.11 Recovery of cleaning detergents in brewing..................... 15.12 Dissolved gas control by membrane technology.............. 15.13 Future trends..................................................................... 15.14 References.........................................................................
430
1 6 Methods for purification of dairy nutraceuticals.................... C. J. Fee, J. M. Billakanti and S. M. Saufi, University of Canterbury, New Zealand 16.1 Introduction ...................................................................... 16.2 Components of acidic whey protein ................................ 16.3 Purification technologies for acidic whey proteins . ........ 16.4 Basic proteins in the dairy nutraceutical industry............. 16.5 Purification technologies for basic whey proteins in the dairy nutraceutical industry............................................... 16.6 Immunoglobulins in the dairy nutraceutical industry....... 16.7 Purification technologies for immunoglobulins in the dairy nutraceutical industry............................................... 16.8 Future trends..................................................................... 16.9 References ........................................................................ 17 Methods of concentration and purification of omega-3 fatty acids . ......................................................................................... S. P. J. Namal Senanayake, Danisco USA, Inc., USA 17.1 Introduction....................................................................... 17.2 Urea adduction in the concentration and purification of omega-3 fatty acids........................................................... 17.3 Chromatographic methods for the concentration and purification of omega-3 fatty acids................................... © Woodhead Publishing Limited, 2010
430 431 432 433 434 435 436 437 441 443 446 446 447 448 450 450 451 454 462 463 470 471 473 474 483 483 484 486
Contents xi 17.4 Low-temperature fractional crystallization for the concentration and purification of omega-3 fatty acids...... 17.5 Supercritical-fluid extraction for the concentration and purification of omega-3 fatty acids................................... 17.6 Distillation methods for the concentration and purification of omega-3 fatty acids................................... 17.7 Enzymatic methods for the concentration and purification of omega-3 fatty acids................................... 17.8 Integrated methods for the concentration and purification of omega-3 fatty acids................................... 17.9 Conclusions....................................................................... 17.10 References......................................................................... 18 Extraction of natural antioxidants from plant foods . ........... E. Conde, A. Moure, H. Domínguez and J. C. Parajó, University of Vigo, Spain 18.1 Introduction....................................................................... 18.2 Antioxidant activity in food systems................................ 18.3 Natural compounds with antioxidant activity and major sources............................................................................... 18.4 Biological activities of natural antioxidants..................... 18.5 Extraction of natural antioxidants from plant foods and residues.............................................................................. 18.6 Integration of extraction processes and purification......... 18.7 Future trends..................................................................... 18.8 Sources of further information and advice....................... 18.9 Acknowledgements........................................................... 18.10 References......................................................................... 19 Fractionation of egg proteins and peptides for nutraceutical applications.......................................................... B. P. Chay Pak Ting, Y. Pouliot and S. F. Gauthier, Laval University, Canada and Y. Mine, University of Guelph, Canada 19.1 Introduction....................................................................... 19.2 Composition and physicochemical characteristics of egg proteins and applications in the nutraceutical industry.............................................................................. 19.3 Biological activities of egg proteins and peptides and applications in the nutraceutical industry......................... 19.4 Available technologies for the fractionation of egg proteins and peptides, and applications in the nutraceutical industry........................................................ 19.5 Conclusion and perspectives............................................. 19.6 References.........................................................................
© Woodhead Publishing Limited, 2010
488 490 492 495 498 501 502 506 506 507 511 521 526 556 567 567 568 568 595
595 597 601 605 612 613
xii Contents 2 0 Supercritical-fluid extraction of lycopene from tomatoes...... J. Shi and S. Jun Xue, Agriculture and Agri-Food Canada, Canada, Y. Jiang, The Chinese Academy of Sciences, China and X. Ye, Zhejiang University, China 20.1 Introduction ...................................................................... 20.2 Supercritical-fluid extraction (SFE) of lycopene.............. 20.3 Factors affecting lycopene yield....................................... 20.4 Effects of pressure and temperature on the antioxidant activity of lycopene........................................................... 20.5 Effect of co-solvent and modifiers in lycopene extraction........................................................................... 20.6 Solubility of lycopene in supercritical fluids.................... 20.7 Conclusion and future trends ........................................... 20.8 References.........................................................................
619
619 622 623 628 631 634 639 640
Index..................................................................................................... 647
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xiii
Contributor contact details
Chapter 2
(* = main contact) Editor S. S. H. Rizvi Department of Food Science Cornell University 114B Stocking Hall Ithaca, NY 14853-7201 USA E-mail: [email protected]
Chapter 1 Prof. Dr Ž. Knez*, Prof. Dr M. Škerget and M. Knez Hrnčič Faculty of Chemistry and Chemical Engineering University of Maribor Slomškov trg 15 2000 Maribor Slovenia
C. Turner* and M. Waldebäck Department of Physical and Analytical Chemistry Uppsala University P.O. Box 599 751 24 Uppsala Sweden E-mail: [email protected]
and C. Turner Department of Chemistry Lund University P.O. Box 124 221 00 Lund Sweden E-mail: [email protected]
E-mail: [email protected]
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xiv Contributor contact details Chapter 3 E. Vorobiev* Groupe Technologies AgroIndustriels EA 4297 Département Génie des Procédés Université de Technologie de Compiègne B.P. 20529 60205 Compiègne France E-mail: [email protected]
F. Chemat UMR 408, Sécurité et Qualité des Produits d’Origine Végétale Université d’Avignon et des Pays de Vaucluse INRA 84000 Avignon France E-mail: [email protected]
Chapter 4 Marcel Ottens* and Sreekanth Chilamkurthi Delft University of Technology Department of Biotechnology Julianalaan 67 2628 BC Delft The Netherlands E-mail: [email protected]
Chapter 5 B. W. Woonton* CSIRO Food and Nutritional Sciences 671 Sneydes Road (Private Bag 16) Werribee Victoria 3030 Australia E-mail: [email protected]
G. W. Smithers Food Industry Consultant P.O. Box 158, Highett Melbourne Victoria 3190 Australia E-mail: [email protected]
Chapter 6 M. Pinelo and A. S. Meyer* Department of Chemical and Biochemical Engineering Center for BioProcess Engineering Building 229 Technical University of Denmark DK-2800 Kgs. Lyngby Denmark E-mail: [email protected]
G. Jonsson CAPEC Center for BioProcess Engineering Building 229 Technical University of Denmark DK-2800 Kgs. Lyngby Denmark Chapter 7 L. Bazinet*, A. Doyen and C. Roblet Institute of Nutraceuticals and Functional Foods (INAF) and Dairy Research Centre (STELA) Department of Food Sciences and Nutrition 2425 rue de l’agriculture, Pavillon Paul-Comtois Laval University Québec, QC, Canada G1V 0A6 E-mail: [email protected]
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Contributor contact details xv Chapter 8
Chapter 12
S. Sahin Department of Food Engineering Middle East Technical University 06531, Ankara Turkey
Dr G. Gésan-Guiziou UMR1253 Science et Technologie du Lait et de l’œuf INRA – Agrocampus Ouest 65 rue de Saint Brieuc 35042 Rennes cedex France
E-mail: [email protected]
E-mail: genevieve.gesan-guiziou@ rennes.inra.fr
Chapter 9 Enrico Drioli* and Alfredo Cassano Institute on Membrane Technology, ITM-CNR c/o University of Calabria via P. Bucci, 17/C I-87030 Rende (CS) Italy E-mail: [email protected];e.drioli@ unical.it; [email protected]
Chapter 10
Chapter 13 Prof. G. Vatai Corvinus University of Budapest Faculty of Food Science Department of Food Engineering 1114 Budapest Menesi út 44 Hungary E-mail: [email protected]
Dr Paula Jauregi* and Dr Maria Dermiki Department of Food and Nutritional Sciences The University of Reading Whiteknights, P.O. Box 226 Reading RG6 6AP UK E-mail: [email protected]
Chapter 14 M. A. Williams Anderson International Corp. 6200 Harvard Avenue Cleveland, OH 44105 USA E-mail: [email protected]
Chapter 11 Dr M.-P. Belleville*, Dr D. Paolucci-Jeanjean and Prof. G. M. Rios Institut Européen des Membranes UMR 5635 CC 047 – UM2 Place E. Bataillon F 34095 Montpellier cedex 05 France E-mail: marie-pierre.belleville@iemm. univ-montp2.fr
Chapter 15 G. J. Freeman Campden BRI Centenary Hall Coopers Hill Road Nutfield Redhill Surrey RH1 4HY UK E-mail: [email protected]
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xvi Contributor contact details Chapter 16
Chapter 19
C. J. Fee*, J. M. Billakanti and S. M. Saufi Biomolecular Interaction Centre Department of Chemical and Process Engineering University of Canterbury Private Bag 4800 Christchurch 8020 New Zealand
Bertrand P. Chay Pak Ting, Yves Pouliot*, Sylvie F. Gauthier Département de Sciences des Aliments et de Nutrition and Institute of Nutraceuticals and Functional Foods (INAF) 2425 rue de l’agriculture, Pavillon Paul-Comtois Laval University Québec, QC Canada G1V 0A6
E-mail: [email protected]
E-mail: [email protected]
Chapter 17 S. P. J. Namal Senanayake Danisco USA, Inc. Four New Century Parkway New Century, KS 66031 USA E-mail: [email protected]
Chapter 18 E. Conde, A. Moure, H. Domínguez* and J. C. Parajó Department of Chemical Engineering Faculty of Sciences University of Vigo Campus Ourense Spain E-mail: [email protected]
Y. Mine Department of Food Science University of Guelph Guelph, ON Canada N1G 2W1 Chapter 20 J. Shi* and S. Jun Xue Guelph Food Research Center Agriculture and Agri-Food Canada Ontario Canada N1G 5C9 E-mail: [email protected]
Y. Jiang South China Botanical Garden The Chinese Academy of Sciences Guangzhou 510650 China X. Ye Department of Food Science and Nutrition School of Biosystems Engineering and Food Science Zhejiang University Zhejiang 310029 China
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xvii
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xxiv Woodhead Publishing Series in Food Science, Technology and Nutrition 178 New technologies in aquaculture: improving production efficiency, quality and environmental management Edited by G. Burnell and G. Allan 179 More baking problems solved S. P. Cauvain and L. S. Young 180 Soft drink and fruit juice problems solved P. Ashurst and R. Hargitt 181 Biofilms in the food and beverage industries Edited by P. M. Fratamico, B. A. Annous and N. W. Gunther 182 Dairy-derived ingredients: food and neutraceutical uses Edited by M. Corredig 183 Handbook of waste management and co-product recovery in food processing Volume 2 Edited by K. W. Waldron 184 Innovations in food labelling Edited by J. Albert 185 Delivering performance in food supply chains Edited by C. Mena and G. Stevens 186 Chemical deterioration and physical instability of food and beverages Edited by L. H. Skibsted, J. Risbo and M. L. Andersen 187 Managing wine quality Volume 1: viticulture and wine quality Edited by A.G. Reynolds 188 Improving the safety and quality of milk Volume 1: milk production and processing Edited by M. Griffiths 189 Improving the safety and quality of milk Volume 2: improving quality in milk products Edited by M. Griffiths 190 Cereal grains: assessing and managing quality Edited by C. Wrigley and I. Batey 191 Sensory analysis for food and beverage quality control: a practical guide Edited by D. Kilcast 192 Managing wine quality Volume 2: oenology and wine quality Edited by A. G. Reynolds 193 Winemaking problems solved Edited by C. E. Butzke 194 Environmental assessment and management in the food industry Edited by U. Sonesson, J. Berlin and F. Ziegler 195 Consumer-driven innovation in food and personal care products Edited by S. R. Jaeger and H. MacFie 196 Tracing pathogens in the food chain Edited by S. Brul, P. M. Fratamico and T. A. McMeekin 197 Case studies in novel food processing technologies: innovations in processing, packaging, and predictive modelling Edited by C. Doona, K Kustin and F. Feeherry 198 Freeze-drying of pharmaceutical and food products T-C. Hua, B-L. Liu and H. Zhang 199 Oxidation in foods and beverages and antioxidant applications Volume 1: understanding mechanisms of oxidation and antioxidant activity Edited by E. A. Decker, R. J. Elias and D. J. McClements 200 Oxidation in foods and beverages and antioxidant applications
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Volume 2: management in different industry sectors Edited by E. A. Decker, R. J. Elias and D. J. McClements Protective cultures, antimicrobial metabolites and bacteriophages for food and beverage biopreservation Edited by C. Lacroix Separation, extraction and concentration processes in the food, beverage and nutraceutical industries Edited by S. S. H. Rizvi Determining mycotoxins and mycotoxigenic fungi in food and feed Edited by S. De Saeger Developing children’s food products Edited by D. Kilcast and F. Angus Functional foods: concept to profit Second edition Edited by M. Saarela Postharvest biology and technology of tropical and subtropical fruits Volume 1 Edited by E. M. Yahia Postharvest biology and technology of tropical and subtropical fruits Volume 2 Edited by E. M. Yahia Postharvest biology and technology of tropical and subtropical fruits Volume 3 Edited by E. M. Yahia Postharvest biology and technology of tropical and subtropical fruits Volume 4 Edited by E. M. Yahia Food and beverage stability and shelf-life Edited by D. Kilcast and P. Subramaniam Processed Meats: improving safety, nutrition and quality Edited by J. P. Kerry and J. F. Kerry Food chain integrity: a holistic approach to food traceability, authenticity, safety and bioterrorism prevention Edited by J. Hoorfar, K. Jordan, F. Butler and R. Prugger Improving the safety and quality of eggs and egg products Volume 1 Edited by Y. Nys, M. Bain and F. Van Immerseel Improving the safety and quality of eggs and egg products Volume 2 Edited by Y. Nys, M. Bain and F. Van Immerseel Feed and fodder contamination: effects on livestock and food safety Edited by J. Fink-Gremmels Hygiene in the design, construction and renovation of food processing factories Edited by H. L. M. Lelieveld and J. Holah Technology of biscuits, crackers and cookies Fourth edition Edited by D. Manley Nanotechnology in the food, beverage and nutraceutical industries Edited by Q. Huang Rice quality K. R. Bhattacharya Meat, poultry and seafood packaging Edited by J. P. Kerry Reducing saturated fats in foods Edited by G. Talbot Handbook of food proteins Edited by G. O. Phillips and P. A. Williams
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xxvi
xxvii
Preface
Separation, extraction and concentration of desirable components from their natural matrices are essential unit operations in the preparation of key ingredients for use in the food, pharmaceutical and chemical industries. These processes often account for somewhere between 50% and 70% of the product cost and are facing new challenges. Ever-tightening regulations on the use of organic solvents, environmental issues and process safety have accelerated the development of a variety of new technologies which are clean and efficient, and do not cause degradation of the products; these technologies therefore enhance the productivity and global competitiveness of the industries concerned. This book aims to provide a comprehensive overview of the most important technologies of interest for the production of high-value compounds. Based on the multidisciplinary expertise of 45 contributors from institutions with strong programs in separation, extraction and concentration processes, the book is organized in 20 peer-reviewed chapters, divided into two parts. Part I describes the latest advances in separation, extraction and concentration techniques, including supercritical fluid extraction, process chromatography and membrane technologies. It also reviews emerging techniques of particular interest, such as pervaporation and pressurized liquid extraction. Part II then focuses on advances in separation technologies and their applications in various sectors of the food, beverage and nutraceutical industries. Areas covered include dairy and egg processing, oilseed extraction and brewing. This part of the book discusses the characteristics of different foods and fluids, how food constituents are affected by separation processes and how separation processes can be designed and operated to optimize end product quality. This volume collectively provides valuable and timely information on the latest developments in the field. With its team of experienced international contributors, Separation, extraction and concentration processes in the food, beverage and nutraceutical industries is an important reference source for experienced professionals concerned with the development and optimization of these processes. It © Woodhead Publishing Limited, 2010
xxviii Preface is hoped that newcomers to this exciting and emerging field will also find valuable information in this book. I wish to thank the authors for their patience and support during the review and preparation of the manuscripts. I also wish to thank Sarah Whitworth of Woodhead Publishing who gave useful advice during the initial planning stages of the book. Syed S. H. Rizvi
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1
Part I Developments in food and nutraceutical separation, extraction and concentration techniques
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2 Separation, extraction and concentration processes
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Principles of supercritical fluid extraction and applications 3
1 Principles of supercritical fluid extraction and applications in the food, beverage and nutraceutical industries Ž. Knez, M. Škerget and M. Knez Hrn�i�, University of Maribor, Slovenia Abstract: The thermodynamic fundamentals of supercritical fluid extraction (SFE) are described and the environmental, health and safety benefits of using supercritical fluids are explored. Several hundred industrial-scale SFE plants are in operation worldwide for extraction of plant materials, such as hop constituents, decaffeination of tea and coffee, and separation of lecithin from oil, all high-pressure processes. Smaller industrial units are used for extraction of spices in the food industry and for natural substances used in cosmetics. The design of such an extraction plant is described. The unique thermodynamic and fluid dynamic properties of dense gases are also applied in integrated extractions and in in situ formulations, such as impregnation of solid particles, formation of solid powder emulsions, and particle coatings. Key words: extraction, supercritical fluids, dense gases, high pressure, thermodynamics, food industry.
1.1 Introduction The design of new products with special characteristics or of new processes that are environmentally friendly and have an impact on sustainability, present a great challenge to chemical engineers. Within the human environment, pressures range from 0.25 bar at the top of the highest mountain, up to 1000 bar at the bottom of the deepest ocean. Because human beings live on the surface of the globe, the first technologies for the production of various substances took place at atmospheric pressure. In the early 20th century, demand for new products like ammonia shifted the technological processes towards high pressure. Industrial high-pressure processes operate at ranges from about 50 bar (in particle formation processes) to over 200 © Woodhead Publishing Limited, 2010
4 Separation, extraction and concentration processes kbar (conversion of graphite to diamonds). High pressure is a relatively new tool and in several processes it has resulted in completely new products with special characteristics. Many of these new processes are environmentally friendly, low cost and sustainable. The advantages of using supercritical fluids (SCF) as solvents in chemical synthesis offer environmental benefits, health and safety benefits and chemical benefits (Jessop and Leitner, 1999). The environmental benefits of most SCFs in industrial processes result from their replacement of far more environmentally damaging conventional organic solvents. The low energy consumption of the process is a further environmental benefit. Health and safety benefits include the fact that the most important SCFs (SC CO2 and SC H2O) are non-carcinogenic, non-toxic, non-mutagenic, non-flammable and thermodynamically stable. One of the major process benefits is derived from the thermophysical properties of SCFs: high diffusivity, low viscosity and the density and dielectric constant of SCF, which can be fine tuned by changes in operating pressure and/or temperature. The motivation for using high pressure in a wide range of technologies and processes is based on chemical, physicochemical, physicobiochemical, physicohydrodynamic and physicohydraulic effects (Bertuco and Vetter, 2001). The extraction of hop constituents and the decaffeination of tea and coffee are the largest scale processes and are mostly performed on an industrial scale. Several industrial plants also extract spices for the food industry and natural substances for use in cosmetics. The advantages of using supercritical fluids for the isolation of natural products have been well described (Marr and Gamse, 2000) and include solvent-free products, low temperature, and no byproducts. One of the most important advantages of using supercritical fluids is the selective extraction of components or the fractionation of total extracts, which is made possible by use of different gases for isolation or fractionation of components and/or by changing the process parameters. A limitation on the further application of high-pressure technology for obtaining extracts is the relatively high price of the product compared with that of those produced conventionally. The legal restrictions on solvent residues and solvents (in products for human use) and the isolation or fractionation of special components from total extracts, in combination with different formulation (controlled release for example) and sterilisation processes, encourage the use of dense gases in extraction applications. There are fewer industrial units involved in using supercritical fluids for the separation of components from liquid mixtures. Some laboratoryscale studies involve extraction systems using liquid/supercritical fluid. Some data on binary systems liquid/SCF has been produced, but there are fewer data on liquid/liquid/supercritical fluid systems. As in all extraction processes, including the supercritical extraction of solid and liquid mixtures, the solubility of a single component or a mixture of components in SCF is the basic data for the design of separation processes. The components
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Principles of supercritical fluid extraction and applications 5 or mixture of compounds to be extracted must be soluble in SCF/dense gas. As is known from thermodynamics, the solubility of compounds in SCF/ dense gases depends on the density of the SCF/dense gas, which depends on the pressure and temperature of the SCF. Another very important parameter influencing the solubility of compounds in SCF is its dielectric constant, which is influenced by the temperature and/or pressure of the SCF. The general flow sheet of the extraction process is presented in Fig. 1.1 and some industrial scale units are shown in Figs 1.2 to 1.4. In one extraction stage, the solubility of a compound or mixture of compounds has to be high whereas in another stage, the solubility of a compound in SCF has to be low. Therefore, the phase equilibrium data is the most important factor in the design of the operating pressure and temperature of SCF in an extraction plant. Thus, the theoretical amount of SCF necessary for separating compounds from a solid or a liquid mixture may be calculated. The design of process parameters has a very important influence on the investment costs of high pressure plants and subsequently, for the economics of the process. In addition to the solubility data of solute in SCF, mass transfer also has a large influence on the economics of the extraction process. Mass transfer models usually describe extraction yield versus extraction time, but a better presentation for the design of extraction apparatuses is yield versus mass of SC solvent/mass of solid material (S/F). Cascade operation is used in industrial-scale plants to increase the economy of the solid–SC solvent extraction process. To date, there has been some research on the continuous operation of plants for the extraction of solids with SC solvent (Eggers, 1996), but currently, no application for
Extractor
PEXT
PS
TEXT
TS
Solvent tank
Extract
Fig. 1.1 General flow sheet of SCF extraction plant. Subscripts EXT and S represent the extractor and separator respectively. © Woodhead Publishing Limited, 2010
6 Separation, extraction and concentration processes
Fig. 1.2 High-pressure extraction unit (700 bar). (Photo: courtesy of Uhde HPT, Hagen, Germany.)
Fig. 1.3 High-pressure extractor solids during manufacturing. (Photo: courtesy of Uhde HPT, Hagen, Germany.)
the continuous feed of solids has been applied on an industrial scale. In industrial scale operations, extractors are usually combined in series. By means of the cyclic operation of a battery of extractors, a quasi-continuous solid flow may be achieved. Such a mode of operation results in extremely high extraction yields because pure solvent is in contact with pre-extracted material, thus loading the solvent to maximum capacity.
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Principles of supercritical fluid extraction and applications 7
Fig. 1.4 High-pressure extraction unit: closure of extractor. (Photo: courtesy of Uhde HPT, Hagen, Germany.)
1
2
3
Solvent tank
4
Extract
Fig. 1.5 Cascade of extractors for extraction of solid materials.
One of the major advantages of SFE processes is the fractionation of extracts. Multi-step separation may be performed by use of several separators by decreasing the solvent power. Decreased solvent power may be achieved by various methods, as described in sections 1.3 and 1.4.
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8 Separation, extraction and concentration processes
Extractor
PEXT
PS1
PS2
PS3
TEXT
TS1
TS2
TS3
Fraction 1
Fraction 2
Solvent tank
Fraction 3
Fig. 1.6 Scheme of multi-step separation.
1.2 Thermodynamic fundamentals Extraction with dense gases is, as with solvent extraction and leaching, a separation process based on the solubility of compounds in phases present in the system. The driving force for mass and heat transfer is the difference from the equilibrium state under given conditions. High-pressure phase behaviour near critical points can be complex, even in simple binary mixtures, especially when the components differ in molecular size, shape, structure and polarity. Complex phase behaviour can be readily interpreted by P–T and P–x projections of P–T–x space diagrams. By using the phase rule (equation [1.1]), the geometrical limits for presentation of multiphase regions in the phase diagram are determined.
F = c + 2 – p
[1.1]
where F is the number of independent variables, c is the number of components and p is the number of phases. The phase equilibrium of compounds (1, 2, 3, ..., N) distributed in phases (a, b, g, ..., p) present in the system can be defined in terms of the fugacity (f) by the following equation:
fia = fib = fig = …… = fip
[1.2]
where i = 1, 2, 3, ..., N. Only binary systems will be discussed in detail in this chapter. The phase equilibria of ternary and multicomponent mixtures is discussed by Brunner (1994) and McHugh and Krukonis (1986), and Sadus (1992). 1.2.1 Solid–supercritical fluid equilibrium Phase diagrams In solid–SCF systems where the normal melting temperature of the solid is higher than the critical temperature (Tc) of the SCF, two possible types of © Woodhead Publishing Limited, 2010
Principles of supercritical fluid extraction and applications 9 phase behaviour exist. The simplest is presented in Fig. 1.7a and is typical of mixtures in which the components are chemically similar. The critical mixture curve runs continuously between the critical points of both components of the mixture. The solid–liquid–gas (SLG) line is continuous and begins at the normal melting point of the heavy component, moves toward lower temperatures as the pressure is increased, and ends at a temperature usually well below the critical temperature of the lighter component. The melting point of the pure solid normally increases with an increase in the hydrostatic pressure. However, in the presence of dense gas, the melting point of the solid decreases as the pressure increases owing to the increasing solubility of gas in the solid. The second type of solid–SCF phase behaviour (Fig. 1.7b) is typical for systems in which the solid and the SCF differ considerably in molecular size, shape and/or polarity and can be interpreted as type III fluid-phase behaviour (de Loos, 2006) according to the classification of van Konynenburg and Scott (1980). In this type of system, the light gas is not very soluble in the heavy liquid, even at high pressures. Therefore, the melting-point depression of the solid is relatively small. The SLG curve is no longer continuous; three phase SLG equilibria are presented by two branches of the SLG line in P–T diagram. The high-temperature branch of the SLG line starts at the normal melting point of the solid and intersects with the critical-mixture curve at the upper critical end point (UCEP). The low-temperature branch of the SLG line intersects with the critical-mixture curve at the lower critical end point (LCEP). At these two points, the liquid and gas phases merge into a single fluid phase in the presence of excess solid. Only solid–gas equilibria exist between these two branches of the SLG line. Possible phase behaviour for type III systems with interference of the solid phase is presented in detail by de Loos (2006). The course of the high-temperature branch of the three-phase SLG line of a binary system depends on the solubility of the gas in the liquid phase. SL
UCEP
1
C2
LG
C1
SLG
2
Pressure (P)
Pressure (P)
SL C1
LG
LG
C2
LCEP SLG
2
1 SLG
TP
TP
Temperature (T) (a)
Temperature (T) (b)
Fig. 1.7 Solid–SCF equilibrium: P–T projection of phase diagram for similar (a) and asymmetrical (b) binary systems. C, critical point; TP, triple point; L, liquid; G, gas; S, solid; UCEP, upper critical end point; LCEP, lower critical end point. Dashed curves are critical lines, lines denoted as 1 and 2 are the vapour pressure curves of the two components. © Woodhead Publishing Limited, 2010
10 Separation, extraction and concentration processes If a compressed gas is dissolved in the melting of a heavy component, two opposite temperature effects occur as given by equation [1.3] (Arons and Diepen, 1963):
fus Ê ∂T ˆ ˆ Ê dT = Á ˜ dP + Á ∂T ˜ dxA = TDVfus d + Ë ∂x A ¯ P Ë ∂P¯ x D A
2 A
fus
d
A
[1.3]
These two effects are that the increase of hydrostatic pressure increases the RT melting temperature of the heavy component, and Pthe dissolved gas x in the D H x H heavy component decreases the melting temperature. TheDhigher solubility of gas in the melted heavy component results in a larger melting-point depression. Four characteristic shapes of the SLG equilibrium lines in the P–T projection were observed experimentally for asymmetric binary systems of compressed gases and non-volatile components (Arons and Diepen, 1963; Loos, 2006; Tuminello et al., 1995; Weidner et al., 1997) (Fig. 1.8), with: ∑ negative dP/dT slope, where the effect of gas solubility predominates; ∑ positive dP/dT slope, where the effect of pressure predominates; ∑ temperature minimum, where both effects are competing; and ∑ temperature maximum and a temperature minimum. The last type of SLG line is a very rare phenomenon and can be explained by the higher solubility of the supercritical gas in the solid phase than in the liquid phase (de Loos, 2006). A SLG curve with a temperature maximum and a temperature minimum was reported for the systems CO2 + polyethylene glycol (Weidner et al., 1997), CO2 + tripalmitin (O’Connell et al., 2003), and
LG
UCEP
Pressure (P)
SLG 2
SLG 1 C2 SLG 4
SLG 3
1
TP
Temperature (T)
Fig. 1.8 Characteristic shapes of the SLG equilibrium lines.
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Principles of supercritical fluid extraction and applications 11 CO2-tristearin (O’Connell et al., 2003). However, in most studies of these equilibria, it is assumed that the supercritical gas is insoluble in the solid phase of the non-volatile component (de Loos, 2006). The course of the SLG line is dependent on the gas and the chemical structure of the compound, i.e. the type and position of the functional groups. As an example, the melting point of vitamin K3 under the pressure of various gases is presented in Fig. 1.9 (Knez and Škerget, 2001) and the various paths of the SLG line can be observed. In the presence of CO2 and dimethyl ether, the negative slope dP/dT can be observed and the meltingpoint depression of vitamin K3 is highest under the pressure of dimethyl ether. The melting-point depression of vitamin K3 is less pronounced in the presence of propane, the SLG curve having a minimum at 94.9 °C and 39 bar. Under the pressure of inert gas (nitrogen and argon), the SLG curve has a positive dP/dT slope owing to the low solubility of gas in vitamin K3. Another example, which illustrates that isomers may have a different type of SLG line in the presence of a specific gas, is the vanillin–gas system. In Fig. 1.10, SLG phase lines for binary systems of vanillin (V) and o-vanillin (o-V) with fluorinated hydrocarbons (R23, R134a, R236fa) and CO2 are presented. For vanillin with –OH group in the para position, the meltingpoint depression in CO2 and fluorinated hydrocarbons is generally lower as for vanillin with –OH group in the ortho position. Thermodynamic modelling For a binary solid–SCF two-phase system at equilibrium, the fugacity of the solute in the solid phase is equal to that in the supercritical phase:
fiS (P, T , x ) = fiG (P, T , y)
[1.4]
350
Propane Dimethyl ether
300
CO2 250
N2
P (bar)
Ar 200 150 100 50 0 50
70
90
T (°C)
110
130
150
Fig. 1.9 SLG phase equilibria for binary system K3 – gas (CO2, propane, dimethyl ether, argon or nitrogen) (Knez and Škerget, 2001).
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12 Separation, extraction and concentration processes 400 V - R23
CHO
350
V - R134a
300
V - R236fa OH
V - CO2 P (bar)
250
OCH3
o-V - R23 CHO
o-V - R134a
200
OH
o-V - R236fa 150
OCH3
o-V - CO2
100 50 0
0
10
20
30
40 50 T (°C)
60
70
80
90
Fig. 1.10 SLG phase lines for binary systems of vanillin (V) and o-vanillin (o-V) with dense gases.
By denoting the gas as component 1 and the non-volatile compound as component 2, and assuming that the solubility of the gas in the solid phase is negligible (x2 = 1), the fugacity of the solid in the solid phase is equal to the fugacity of pure solid (Prausnitz et al., 1986):
f2S (P, T , x ) = f2S,pure (P, T )
[1.5]
The fugacity of a pure compound is:
RT ln
f2S,pure (P,T ) = P
Ú0
RT ln
f2S,pure (P,T ) = P
Ú0
P
P2S
RT ˆ Ê ÁË v2 – P ˜¯ dP RT ˆ Ê ÁË v2 – P ˜¯ dP +
[1.6] P
ÚP
S 2
Ê S RT ˆ ÁË v2 – P ˜¯ dP [1.7]
where P2s is the sublimation pressure at the system temperature and v2S is the molar volume of the pure solid. The first term on the right side is the fugacity of the saturated vapour, which is equal to the fugacity of the saturated solid phase. The second term is the correction owing to the compression of the solid phase to pressure P.
RT ln
f2S,pure (P,T ) fS = RT ln 2S + P P2
P
ÚP
S 2
v2S dP S P – RT ln P2 P
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[1.8]
Principles of supercritical fluid extraction and applications 13 By rearranging and inserting the expression for the fugacity coefficient fS j 2S = 2S : P2
Ê f2S,pure (P, T ) = P2Sj 2S exp Á 1 Ë RT
P
ÚP
S 2
ˆ v2S dP˜ ¯
[1.9]
where j2S is the fugacity coefficient at T and P2S. The fugacity of the gas phase is:
f2G (P, T , y) = j 2G y2 P
[1.10]
j2G
is the fugacity coefficient of solid component 2 in the mixture. where By inserting equations [1.9] and [1.10] into [1.5] and [1.4], the expression for solubility of solids in the gas phase is obtained:
P Ê ˆ P2Sj 2S expÁ 1 Ú S v2S dP˜ Ë RT P2 ¯ y2 = G Pj 2
[1.11]
Equation [1.11] can also be expressed in the form:
y2 =
P2S E P
[1.12]
where
P Ê ˆ j 2S expÁ 1 Ú S v2S dP˜ Ë RT P2 ¯ E∫ G j2
[1.13]
The enhancement factor E is the correction of the ideal-gas expression that is valid only at low pressures and contains three terms: j2S, which takes into account the non-ideality of the pure saturated vapour. For low sublimation pressure of the solid P2S, j2S almost equals unity. ∑ an exponential term called the Poynting correction, which gives the effect of pressure on the fugacity of the pure solid. It is small at low pressures but may become larger at high pressures or at low temperatures. ∑ j2G, the gas-phase fugacity coefficient in the high-pressure gas mixture. This term is the most important, because it is much lower than 1 and can produce very large enhancement factors (103 or higher). ∑
Assuming that the molar volume of the pure solid (v 2S) at the system temperature is pressure independent, the Poyning correction takes the simple form:
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14 Separation, extraction and concentration processes
Ê vS (P – P2S )ˆ j 2S expÁ 2 ˜¯ RT Ë E= G j2
[1.14]
j2G can be calculated from an equation of state using following equation:
lnj i = 1 RT
•
ÚV
ÈÊ ∂P ˆ ˘ – RT ˙ dV – ln PV ÍÁ ˜ nT RT V ˙ ÍÎË ∂ni ¯ T ,V , n j ˚
[1.15]
where ni is the molar number of species i in the mixture, V is the total volume, v is the molar volume, and R is the gas constant. To calculate fugacity coefficients, the equations of state (EOS) are commonly applied in engineering practice because they can be used to predict the thermodynamic properties of fluids and describe the phase behaviour of mixtures over a wide range of temperature and pressure. Detailed reviews on the thermodynamic models applied for predicting phase behaviour and modelling aspects in supercritical fluid mixtures have been presented elsewhere (Anderko, 1990, 2000; Dohrn, 1994; Tuminello et al., 1995). Cubic EOS in combination with mixing rules are currently the most widely used models for the calculation of solubilities of components in SCF (Tables 1.1 and 1.2). Table 1.1 Cubic equations of state General form
Van der Waals (VDW: b = 0, c = 0) Redlich–Kwong (RK: c = 0)
Redlich–Kwong–Soave (RKS: c = 0) Peng–Robinson (PR: c = b)
a P = RT – v – b v(v + b ) + c(v – b )
[1.16]
a P = RT – 2c v–b v
[1.17]
acTc1/2 P = RT – v – b v(v + b )T 1/2
[1.18]
a P = RT – v – b v(v + b )
[1.19]
a P = RT – v – b v(v + b ) + b(v – b )
Where: a = aca
ac = Wa
[1.20]
[1.21]
RT Pc 2
2 c
[1.22]
a = [1 + c(1 – Tr1/2)]2
[1.23]
2
c = A 0 + B 0w + C 0w
[1.24]
RTc Pc
[1.25]
b = Wb
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Principles of supercritical fluid extraction and applications 15 Table 1.2 Constants in cubic equations of state Konstanta
VDW
RK
RKS
PR
Wa Wb A0 B0 C0
0.42188 0.125 – – –
0.42478 0.08664 – – –
0.42747 0.08664 0.48 1.574 –0.176
0.45724 0.0778 0.37464 1.54226 –0.26992
In the equations in Tables 1.1 and 1.2, the parameters a and b reflect the contribution of attractive forces and molecular volume, respectively, and w, the acentric factor, is a measure of the acentric nature of intermolecular forces. In order to extend the use of a pure-fluid EOS to mixtures, it is assumed that the EOS for the mixture is the same as for a hypothetical pure fluid (Prausnitz et al., 1986) and that the characteristic constants a and b are dependent on composition. Usually, van der Waals one-fluid mixing rules with one or two adjustable parameters are used:
a = ∑∑ yi y j aij i j
b = ∑∑ yi y j bij i j
[1.26]
[1.27]
where a accounts for interactions between the species in the mixture and b accounts for the excluded volume of the species of the mixture. The cross coefficients aij and bij are related to the corresponding pure-component parameters by the following combining rules:
aij = aii a jj (1 – kij )
bij = 1 (bii + b jj )(1 – cij ) 2
[1.28] [1.29]
where kij and cij are the binary interaction and size parameters, respectively. When cij is set to zero, the co-volume b is expressed by a linear mixing rule, equation [1.29] reduces to:
b = ∑ yi bi i
[1.30]
In most practical applications, a linear mixing rule is used for the co-volume b (Anderko, 1990). However, many authors use a quadratic mixing rule for b in analogy with the mixing rule for a. In this way, a second binary parameter is introduced, which is usually found to be useful in correlating gas–liquid equilibria in mixtures with components of very different size (Anderko, 1990). Similarly, it usually improves results for solid–supercritical fluid equilibria © Woodhead Publishing Limited, 2010
16 Separation, extraction and concentration processes because of the large size differences in such systems (Anderko, 1990). Generally, kij and cij are of the same order of magnitude (Dohrn, 1994) and both are expected to have an absolute value of much less than 1 (McHugh et al., 1986). They can both be positive or negative. A negative value of kij usually indicates that specific chemical interactions, such as hydrogen bonding, are present in the mixture (McHugh et al., 1986). However, it is less apparent how to interpret a negative value for cij. By calculating partial derivatives of cubic EOS, inserting derivatives in equation [1.15] and after integrating the expression for the fugacity coefficient of the component, i is obtained. The expression for the fugacity coefficient of component i by using the Peng–Robinson (PR) EOS with the above defined van der Waals mixing rules is:
lnj i =
bN Ê Pv ˆ – 1˜ b ÁË RT ¯
Ê N ˆ 2∑y j aij Á b ˜ P(v – b ) j a – ln – – N˜ Á b˜ RT 2 2 RTb Á a ÁË ˜¯
¥l
[1.31]
where:
N
N
N –1 N
k
j
j =1i = j +1
bN = 2∑yk bik – ∑ y 2j b jj – 2 ∑ ∑ yi yi –j
ij
and for a binary system: v + (1 + 2)b y b ln bN = 2yi bij +v 2+y(1 yi2bbii – y 2j b jj – 2yi y j bi j b jj– – 2)
[1.32]
[1.33]
ij in the above equations The critical properties and acentric factors needed can usually be estimated by the use of group contribution methods, molar volume can be determined experimentally e.g. with a pycnometer, whereas the binary interaction and size parameters, kij and cij, respectively, are obtained by fitting the equation of state to measured phase-equilibrium data. The form of the mixing rules that extend the use of EOS developed for pure fluids to mixtures is, as reported (Anderko, 1990; Škerget et al., 2002), more important than the particular P–V–T relationship embodied in EOS. The shortcoming of the van der Waals one-fluid mixing rules is that they are only applicable to mixtures that exhibit relatively moderate solution nonidealities. Further problems, which persist when modelling phase equilibria in SCF mixtures with the use of cubic equations of state where the systems are highly non-ideal owing to the high pressures involved, are described by Prausnitz et al. (1986) and Škerget et al. (2002):
(a) the interaction parameters are often found to be temperature dependent; © Woodhead Publishing Limited, 2010
Principles of supercritical fluid extraction and applications 17 (b) reliable values for the necessary physical parameters are not always available; and (c) the equation does not fit the data equally well at all temperatures and pressures, and especially in the vicinity of the critical point, the deviation from the model is high. There have been a number of attempts to extend the range of EOS. Most equations retain the van der Waals separation of repulsive and attractive terms, but have introduced some modifications to either the attractive or repulsive term, or to both. In this way, a number of empirical and theoretical models have been developed (Sandler and Orbey, 2000, Sadus, 1992), which allow extrapolation and prediction over wide ranges of temperature and pressure and can describe greater degrees of non-ideality. Empirical approaches to eliminating the shortcomings of the van der Waals one-fluid model for a cubic EOS have been done to provide additional composition dependence or density dependence, by adding parameters to the combining rule for the parameter a, generally leaving the mixing rule for the parameter b unchanged (Sadus, 1992, Sandler and Orbey, 2000, Škerget et al., 2002). Some examples are the composition dependent combining rules of Adachi and Sugie (1986), Panagiotopoulos et al. (1986), Sandoval et al. (1989) and Schwartzentruber and Renon (1989 a,b). The density-dependent mixing rules have been revived by Anderko (1990) and Danner and Gupte (1986). A more empirical way was adopted by Luedecke and Prausnitz (1985). However, there are several problems associated with these multi-parameter combining rules which limit their use (Anderko, 1990; Škerget et al., 2002). The common difficulty is that they do not result in the correct treatment of ternary and multicomponent mixtures (de Loos, 2006). An alternative method is based on the combination of the equations of state with activitycoefficient models (Huron and Vidal (1979) and Wong and Sandler mixing rules (Wong and Sandler, 1992), which are usually fairly reliable for the prediction of multicomponent phase equilibria from binary data, except for the most strongly non-ideal systems. In binary three-phase SLG systems at equilibrium, the parameters needed in the PR EOS to model the SLG line are determined by solving the following equations:
f 2G(T, P, y2) = f 2L(T, P, x2)
[1.34]
f 2S (T,
[1.35]
f 1G(T, P, y1) = f 1L(T, P, x1)
P) =
f 2G(T,
P, y2)
[1.36]
Equation [1.35] transforms to equation [1.11] and equations [1.36] and [1.34] become:
y1j1G = x1j1L
[1.37]
y2j2G
[1.38]
=
x2j2L © Woodhead Publishing Limited, 2010
18 Separation, extraction and concentration processes Calculation of the pure solute fugacity The fugacity of a solute in a solid phase cannot be directly calculated by a conventional EOS, but it may be calculated by means of a sub-cooled liquid reference state. An equation developed by Prausnitz et al. (1986) is used to express the fugacity of the sub-cooled liquid at temperature T in terms of measurable thermodynamic properties: ln
f2L,pure DH 2fus Ê Tt,2 ˆ = ÁË T – 1¯ S,pure RT t,2 f2
Dcp,2 Tt,2 R T
Dcp,2 R
Tt,2 T [1.39]
where Tt,2 is the triple-point temperature, and DH 2fus the enthalpy of fusion Ê ˆ for component 2 at temperature T. ˜ – – 1˜ + ln ÁË ¯ ¯ Usually, it is sufficient to consider only the first term in equation [1.39], especially, if T and Tt are not far apart, and to neglect the contribution of cp. Furthermore, because there is little difference between the triple-point temperature and the normal melting temperature, it is common to substitute Tt for its normal melting point Tfus,i. The simplified equation thus obtained, valid at the triple point pressure Pt of the solute, is: f2L,pure DH 2fus Ê Tfus,2 = –1 [1.40] ˆ f2S,pure RTfus,2 Ê T ÁË ˜¯ To take into account the effect of pressure, the fugacities may be written as: ln
f2S,pure (P, T ) = f2S,pure (Pt , T ) exp
Ê 1 RT
ÊÊ 1 f2L,pure (P, T ) = f2L,pure (Pt , T ) expÁ Ë RT
P
ÚP v2S d P t
ˆ
P
ÚP v2L d P˜¯
[1.42]
t
P
ˆ ˜¯ (v – v ) dP˘ ˙˚
and the fugacity ratio is:
Ê ÁË f2S,pure (P, T ) f2S,pure (Pt , T ) = L,p exp È 1 L,pure ÍÎRT f2 (P, T ) f2 (P , T )
[1.41]
ÚP
t
By inserting equation [1.40] pureinto equation [1.43]:
S 2
L 2
[1.43]
t
f2S,pure (P, T )
and by integrating:
=
È f2L,pure (P, T )exp Í
1 RT Î
P
ÚP (v2S – v2L )dP + t
D
fus 2
Ê Á1 – fus,2 Ë
fus,2 ˆ ˘
˜¯ ˙ ˚
[1.44]
P
È(vS – v2L )(P – Pt f2S,pure (P, T ) = f2L,pure (P, T ) exp Í 2 RT Î –
)
© Woodhead Publishing Limited, 2010
+
T H RT T DH 2fus Ê Tfus,2 ˆ ˘ 1 RTfus,2 ÁË T ˜¯ ˙˚ –
[1.45]
Principles of supercritical fluid extraction and applications 19 1.2.2 Liquid–SCF equilibrium Phase diagrams According to the classification proposed by van Konynenburg and Scott (1980), six main types of fluid-phase behaviour in binary mixtures are distinguished by their critical properties. The corresponding P–T projections are shown in Fig. 1.11. Similar behaviour is observed for type I and type II, where a continuous LG critical line, which connects the critical points of the pure components, is observed. However, the difference between both types is observed because type II liquid mixtures are not miscible in all proportions and exhibit a LLG three phase line at low temperatures. The LLG line ends at the UCEP, where two liquid phases merge into one liquid phase. From this point, the LL critical line rises rapidly to higher pressures. In type III phase behaviour, the LG critical line is not a continuous line connecting critical points of the pure components, instead it has two branches; one going from the critical point of the component with the higher critical temperature, initially in the direction of the critical point of the other component and then rising sharply with pressure, whereas the other line goes from the critical point of the other component to the UCEP. Various sub-classes of type III behaviour have been observed, depending on the course of the main critical line. The critical line starting from C2 with a positive slope, indicates the existence of the so-called gas–gas equilibria Type I
Type II
1 Pressure (P)
C2
LG
C1
LL
Type III
1
2
Type V LG
C2 LL
C2
C1
LG
UCEP
LCEP
UCEP LLG
2
LCEP
1 LLG
UCEP
LL
2
1
C2
LG
C1
LG
LLG
2
LLG
Type VI LG
C1
1
LG
UCEP LLG
Type IV
C2
LG C1
1
2
C2
LG
C1
UCEP LLG
2
LCEP
Temperature (T)
Fig. 1.11 Classification of the phase behaviour of binary fluid systems. C, critical point; L, liquid; G, gas; UCEP, upper critical end point; LCEP, lower critical end point. Dashed curves are critical lines, lines denoted as 1 and 2 are the vapour pressure curves of the two components (Brunner, 1994; Prausnitz et al., 1986).
© Woodhead Publishing Limited, 2010
20 Separation, extraction and concentration processes (Sadus, 1992). The two phases are in equilibrium at a temperature higher than the critical temperature of either pure component. The three-phase LLG line is also observed near the vapour pressure line of the most volatile compound. In type IV phase behaviour, the critical line between the critical points of pure compounds is interrupted by a three-phase LLG line. One branch of the critical line goes from the critical point of the component with the higher critical temperature to the LCEP, whereas the other branch goes, as in type III, from C1 to the UCEP. Type V is similar to type IV except that there is also an LL critical line at low temperatures which ends at another UCEP. For this type of mixture, there are therefore two regions of limited liquid miscibility at lower and higher temperatures and pressures. In type VI there are two critical curves: a continuous LG critical line, which connects the critical points of the pure compounds, and an LL critical curve, which connects the UCEP and LCEP located at either end of threephase LLG line. Thermodynamic modelling of high-pressure vapour–liquid (V–L) equilibria Consideration of binary liquid in equilibrium with a gas phase: the equilibrium equation for a binary liquid–gas system at temperature T and pressure P is:
f 1G(T, P, y1) = f 1L(T, P, x1)
[1.45]
f 2G(T, P, y2) = f 2L(T, P, x2)
[1.46]
where xi is the mole fraction of component i in the liquid phase and yi is the mole fraction of component i in the gas phase. The fugacity in each phase can be written as:
f Li = xi jLi P
[1.47]
f Gi = yi jGi P
[1.48]
where jGi and jLi are the fugacity coefficients of i in the gas and liquid phases, respectively.
y1j1V = x1j1L
[1.49]
y2jV2 = x2 j2L
[1.50]
Fugacity coefficients for each component in the liquid and gas phase can be calculated by using equation [1.15].
© Woodhead Publishing Limited, 2010
Principles of supercritical fluid extraction and applications 21
1.3 Cycle processes for extraction using supercritical fluids 1.3.1 Solvent cycle A typical high-pressure extraction process basically comprises a separation stage for the feedstock (extractor and separator) and a regeneration stage for the solvent. In the separation stage, the components to be extracted become concentrated in the gas and are then precipitated in the separator by applying suitable methods (Fig. 1.12). The gas must be removed from the extract and the raffinate and cleaned for reuse in the extraction process. If a solvent mixture is applied in the extraction process, the composition of the mixture must be adjusted before reuse. The solvent recovery can be performed in various ways and the procedure chosen depends on the nature of substances, the scale of the process unit and the operating conditions (Brunner, 1994). The extract is separated from the solvent by changing the operating conditions in such a way as to reduce the solvent power of the SCF. This can be achieved by various means (Brunner, 1994): ∑
Isenthalpic throttling (expansion) to subcritical conditions (changing pressure and temperature), ∑ Changing the temperature and maintaining supercritical conditions, or cooling down to subcritical conditions, ∑ Employing an additional mass separating agent (absorbing medium, membrane adsorbents), whilst maintaining supercritical conditions for the solvent. Separation stage
Material to be extracted
Regeneration stage
Gas
Solvent–solute mixture
Extractor
Separator
Gas reservoir
Gas Raffinate
Gas
Solvent (compressed gas)
Extract (solute)
Fig. 1.12 Solvent cycle in extraction process using SCF. © Woodhead Publishing Limited, 2010
22 Separation, extraction and concentration processes The appropriate method and conditions in the separator are chosen according to the phase equilibrium data. In general, a change in temperature will not be effective in cleaning the solvent sufficiently for reuse in the extraction process, but a change in temperature can be applied in addition if total regeneration of the solvent is not necessary (Brunner, 1994). The advantage of using an absorbent or adsorbent for separation of the extract and the solvent is that the separation can be performed without significantly reducing the pressure. Operating at constant pressure reduces energy consumption. The disadvantage is that separation of the extract and the absorbent/adsorbent have to be performed subsequently, and this may be difficult and must be taken into account when designing the process and calculating its operating costs. The solvent in the subcritical (liquid) or supercritical state may be driven either by a compressor or by a pump. Solvent cycles with corresponding changes of conditions of state can be best presented in T–S diagrams and are used for calculating the heat balance of each step of the process. Pump process In Fig. 1.13 the temperature–entropy (T–S) diagram is schematically presented with areas of homogeneous liquid, gas, SCF and two-phase liquid–gas region (L–G). The dotted line separates the area of SCF from the gas and liquid state area and presents no phase border. The part of the line on the right of the critical point (CP) lies on the critical isobar (73 bar for CO 2). The extraction process is performed under constant conditions (3 in Fig. 1.13). The SCF phase following the extraction leaves the extraction unit and the dissolved substance is subsequently separated from the solvent by changing the pressure and temperature to reduce the solvent power of the SCF (isenthalpic H6
P5
P4
H5
Temperature (T)
P3 H4 H3 H2 2
H1 L
Pc
SCF
3
P2 P1
CP
G 5
1
6
L–G
4
Entropy (S)
Fig. 1.13 Gas extraction process in T–S diagram: solvent circuit in the pump mode.
© Woodhead Publishing Limited, 2010
Principles of supercritical fluid extraction and applications 23 throttling 3–4). The two-phase region of the solvent is obtained and the substance dissolved in the solvent precipitates to form a separate phase that can be removed. The gaseous part of the solvent is cooled and condensed to a liquid state (6) and then sub-cooled (1). To remove the substances that remain dissolved, the liquid solvent can be evaporated (6–5), subsequently liquefied (5–6) and sub-cooled (1). The sub-cooled liquid is then pumped to the operating pressure (1–2) and heated to the operating temperature (2–3). The area enclosed by the solvent circuit in the T–S diagram represents the thermodynamic work needed for the process of cycling the SCF. The lower the pressure drop needed for separating the extract and the solvent, the lower will be the work required. Compressor process The solvent circuit in the extraction process using a compressor is schematically presented in the T–S diagram in Fig. 1.14. After the extraction process, which is performed under constant conditions (3), the solute is separated from the SCF by isenthalpic throttling into the subcritical region (3–4) and this is followed by evaporation (4–5–1). The dissolved substances precipitate and the solvent, which is in a gaseous state, is then compressed to the pressure of the extraction, shown in Fig. 1.14 as an idealised, isentropic compression (1–2). During compression, the temperature of the SCF increases, therefore the solvent is cooled to the temperature of extraction (2–3). This completes the cycle, which may be repeated. The operating modes in Figs 1.13 and 1.14 are presented as examples. Solvent cycles may vary for different gases and depend upon the mode of H6
P5
H5
P4 P3
Temperature (T)
H4 H3
Pc
SCF
2
H2
H1
P2 P1
3
CP
G 1/5
L
4
L–G
Entropy (S)
Fig. 1.14 Gas extraction process in T–S diagram: solvent circuit in the compressor mode.
© Woodhead Publishing Limited, 2010
24 Separation, extraction and concentration processes operation, e.g. the stages involved in separating the extract from the solvent can be done by pressure decrease, temperature increase, temperature decrease or a combination of temperature and pressure change. The mass and heat losses should also be considered when calculating the balance for practical processes. 1.3.2 Separation of solute in extraction processes using SCF After the solubilisation of SCF soluble substances in the extractor, the dissolved substances must be separated from the solvent–SCF mixture. Several separation processes may be used, but the preferred option is to use processes in which no additional substances are introduced into the system, which have varying solvent powers for supercritical fluid, and where only small changes in the conditions of the extraction process are made. Adding substances to separate solute from the SCF increases the costs of separation, whereas re-compression and cooling conditioning of the SCF to operating extraction conditions, increases the energy costs of the process. One of main advantages of using SCF for extraction and fractionation of substances, is that by a stepwise reduction of solvent power, fractionation of extracts can be obtained. Separation of solute by reduced solvent power Reduction of pressure and temperature increase The mostly widely used process for the separation of solute from solvent, is reduction of the solvent power. A condensed phase is formed and the gaseous phase is separated. The solvent power of supercritical solvents depends upon pressure and temperature which influence the density of the SCF. Generally, the solubility of a solute increases with increasing density and decreases with decreasing density. A decrease in density can be obtained by decreasing pressure and/or increasing temperature. In many industrial operations, the concentration of solute in a solvent is reduced by decreasing the pressure and thus the density, whereas an increase in temperature increases the vapour pressure of the solute. In a low pressure range where the density decreases dramatically with a temperature increase, the concentration of solute in SCF decreases with increasing temperature when the vapour pressures of the substances are relatively low. The separation of a solute by temperature increase is not an efficient method (for example for separation of essential oils from SCF) when the vapour pressure of a solute is high and increases with a rise in temperature. High pressure, where density decreases slightly with an increase in temperature, will lead to a higher concentration of solute with increasing temperature (at constant pressure) in the SCF. Based on these observations, it is evident that the increase in temperature, which decreases the density and solvent power, should provide an efficient method for the separation of a solute with low or moderate vapour pressure within the low-pressure range. On the contrary,
© Woodhead Publishing Limited, 2010
Principles of supercritical fluid extraction and applications 25 for some substances, at certain process conditions, the solubility decreases with increasing pressure. For such systems, separation procedures of solutes by pressure reduction or reduction of the solvent power is not possible. It is evident that the design of efficient processes for the separation of solute from SCF by the reduction of solvent power is possible only when phase equilibrium data are available. After establishing thermodynamic conditions for reduced solubility in a separator, a condensed phase is formed and this is separated from the solvent with lower density in one, or several separators in series (in the last separator from the gaseous phase). Usually, such a series of separators will operate at different pressures and/or temperatures to fractionate the single components of the extract. The solute–solvent system must remain in the separator for a sufficient length of time to approximately achieve the phase equilibrium. Separation by expansion Under certain process conditions, the solute-loaded solvent phase is expanded in the separator where the solvent is a subcritical liquid. In this instance, both the gaseous solvent and the liquid solvent containing the solute are present in the separator, a three-phase mixture is present if the solute is insoluble or less soluble in the liquid and gaseous phases. Usually, the liquid solvent evaporates and the solute is removed from the separator, whereas the gaseous phase is reused. This set of separation sequences is effective for the separation of a solute from solute-loaded SCF, but it is extremely expensive owing to the high energy input required. For efficient separation of the condensed phase from low-density solvent, auxiliary devices are used to increase the separation of the solid or liquid phase from the low-density solvent. These include demisters consisting of wire mesh packing, deflectors and filters, and cyclones. Separation of solute and solvent by a mass separating agent The separation of solute from a solution by a mass separating agent is possible by: ∑ absorption, ∑ adsorption, ∑ use of membranes, or ∑ adding a substance of low solvent power. For separation of a solute from a solvent by absorption, the solvent circuit may operate at an almost constant pressure. The absorbing liquid must dissolve the solute and should not absorb the solvent. Separation of solute from the SC solvent by adsorption can be a very efficient process. This process, like that of absorption, can operate with practically no pressure and/or temperature drop. Therefore, both separation processes have a major influence on the economy of whole extraction process. Membrane separation processes may be efficiently used for the separation
© Woodhead Publishing Limited, 2010
26 Separation, extraction and concentration processes of solute from gas phase, owing to the difference in molecular mass. The pressure drop through the membrane is relatively low, therefore the solvent regeneration costs are low. The solvent power of a SC solvent may be reduced by adding a substance of low solvent power (e.g. adding nitrogen to SC CO2 for several substances). A similar effect may be obtained, when the entrainer is separated from the SC solution (if entrainer was applied in extraction process). For separating the solute, the entrainer has to be removed by adsorption or absorption. The advantage of such a separation process is its ability to operate at an almost constant pressure, thus keeping costs to a minimum.
1.4 Extraction of solids using SCF 1.4.1 Design of extraction plant The fundamentals of the design criteria for SCF extraction of solids in relation to process and equipment are reviewed in this section. For successful engineering and/or design of a SC extraction process, the following parameters should be defined: ∑ ∑
specific basic data, thermodynamic conditions for the operation of extraction and separation process, ∑ mass transfer data for the system, and ∑ energy consumption by means of T–S diagrams. The size of extractors for a certain capacity of plant are defined by specific basic data and thermodynamic conditions. Mass transfer data determines the time solids remain in contact with the SC solvent, so determining the solvent circulating system, including energy (heating and cooling) consumption, and therefore the capacity of the pump/compressor, the surface of heat exchangers and the piping system. Specific data and pretreatment of solids For the design of industrial plant, the following information is essential: ∑ raw material specification, ∑ final product specification, ∑ required plant size, ∑ plant location, sometimes taking into account the prevailing local conditions. The specification of the raw materials influences the quality of the extracts obtained and the overall economy of the process. If the raw material is contaminated, either by waste products or plant protection materials, a specified level of purity in the extract cannot usually be achieved. On the other hand, if the raw material contains a low concentration of substances to be extracted, the economy of the process will be questionable. © Woodhead Publishing Limited, 2010
Principles of supercritical fluid extraction and applications 27 The final product specification is usually based on customer needs, expectations and specifications and has a considerable influence on the costs of the extraction process. The process conditions in the extraction stage should be determined in order to obtain maximum yield and high selectivity for the required substance at minimal separation costs. However, a multistep separation is required for the fractionation of extracts. From a previous study (Gamse and Marr, 2001) it is known that the extractor volume influences the investment costs for an extraction unit (investment costs are a logarithmic function of extractor volume). Location will determine the mechanical construction of an extraction plant [e.g. location in Good Manufacturing Practice (GMP) area or in an earthquake area]. Climate conditions will also influence the design of heating and cooling devices and electrical drives. Pretreatment of solids Substances which should be extracted with SC fluids may be divided in two categories: ∑ ∑
raw materials where only unwanted substances should be removed and the geometry of the raw material is maintained during the process, and raw materials where pretreatment is allowed and extracted substances are the main product.
Examples of the first group of materials are coffee and tea decaffeination, the defatting of some seed plant materials, and the separation of pesticides from rice and ginseng. In the first process, high selectivity for the separation of unwanted compounds is necessary and therefore the process parameters or isolation of substances have to be very precisely selected. These processes are used for high-volume and relatively high-value market products. For the other group of materials, any pretreatment may typically be used to achieve a high yield at a low solvent to feed ratio with a low energy input. Therefore, the raw material should be ground before extraction to increase the bulk density. Some non-ground material could have a bulk density of 150–250 kg m–3, whereas the same material when ground would have a bulk density of 350–500 kg m–3. The ground feed material has an average particle size of 0.4 to 0.8 mm. Smaller particles have a higher specific area and therefore a higher mass flux could be expected. However, the linear solvent velocity through the ground raw material is decreased and consequently a greater drop in pressure is obtained. The linear solvent velocity through the raw material is 0.7 to 0.8 cm s–1. Finely ground material may also cause clogging of sintered plates on extraction vessels or filters and tends to cause channelling. If the bulk density of a ground substance is less than 250 kg m–3, the material should be formed into pellets. Bulk density has a strong influence on the economics of the process. If the bulk density in the extraction vessel is too low, less raw material is loaded and the yield per batch is reduced. Therefore, for a given product capacity, the volume of the extraction vessels should be greater. The moisture content
© Woodhead Publishing Limited, 2010
28 Separation, extraction and concentration processes of raw plant materials is between 8 to 15 wt%. The amount of water in plant material influences the economics of the process. In plant material having a higher water content, the relatively high solubility of water in CO2 means that polar substances are extracted at relatively low pressure. If the water concentration in the raw plant material is too low, the cells may shrink and hinder the mass transfer of substances which are extracted through the cell walls. The ideal moisture content for extraction of substances from plant material should be optimised by laboratory-scale tests. It is known that for the decaffeination of coffee and tea, the water content of green coffee beans should range from 35 to 45 wt% (Zosel, 1981), whereas in the extraction of astaxanthin from algae, it should be as low as possible (water–oil emulsions are formed and the yield of astaxanthin would therefore be very low). In the SC extraction process for the isolation of colorants and antioxidants from raw plant materials, the content of oils, essential oils and waxes has a particularly large influence on the extraction yield. These substances contained in raw plant material may act as entrainers for the extraction of valuable compounds (such phenomena could be observed for extraction of carotenoids from ground paprika, where, owing to the low solubility of carotenoids in pure CO2 even at ultra high pressure (1500 bar), the carotenoids could not be separated quantitatively from plant material. Therefore, in some processes, oils, essential oils and waxes are used as entrainers for extraction of valuable compounds. Thermodynamic data The solubility of an extracted substance in a supercritical solvent is critical to the economics of the extraction process. The highest possible loading of SC solvent should be achieved in the extraction stages of the process, whereas in the separation stage, the solubility of solute in solvent should be as low as possible. In some instances, the solvent power for solutes should be gradually reduced so the fractionated separation of solute from the solvent may be achieved. Details of the thermodynamic properties of the solute in SC solvents are described in section 1.2. Mass transfer Mass transfer data are essential for determining the extraction time and the capacities of pumps, compressors, and heat exchangers. The data on mass transfer should be optimised experimentally with raw plant material having defined humidity, particle size, and particle size distribution at defined optimal extraction pressure and optimal extraction temperature (data obtained from thermodynamic investigations) for a defined SC solvent. Based on mass transfer experiments, the mass of solvent per mass of feed should be determined for the highest possible yield of the substance to be extracted. Figure 1.15 presents the typical extraction curve for the isolation of a substance from solids. The extraction curve for the isolation of substances from solid raw plant material is generally divided into a period of constant
© Woodhead Publishing Limited, 2010
Principles of supercritical fluid extraction and applications 29
Yield (wt%)
Diffusion-controlled mass transfer
Solubility-controlled mass transfer
Mass of solvent/mass of feed (S/F)
Fig. 1.15 Typical extraction curve.
extraction rate and a period of falling extraction rate (Sovova, 2005). Solubility controlled mass transfer prevails in the initial period of the process, whereas at higher S/F (longer extraction time), the mass transfer is controlled by diffusion. The diffusion and hydrodynamics influence the mass transfer rates. Diffusion The substance to be extracted may be located in the cells of the raw material or adsorbed on the surface of a solid matrix. Therefore, the mass transfer depends on the location of the substance. If it is adsorbed on the surface, the mass transfer rates are high and, vice versa, when the solid particles of the substances diffuse through cell walls, the mass transfer rates will be low. In some instances, the substances to be extracted also form complexes which require release by a chemical reaction (usually hydrolysis with water). For substances that do not form a chemical complex, the diffusion may be influenced by: ∑ ∑
reduction of particle size (reduction of diffusion path), and destruction of cells (by swelling or cracking cells, by ultrasound, by milling procedure).
Hydrodynamics Particle properties such as particle size, particle shape and particle size distribution may cause channelling, which reduces the flow of the SC solvent through the material to be extracted. Some materials swell during the extraction process and may reduce the flow of the SC solvent through the extractor. For some substances, the direction of flow of the SC solvent is important. In industrial scale processes, the flow will usually be from bottom to top, whereas under some process conditions, the opposite direction of flow
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30 Separation, extraction and concentration processes will give higher yields. Mechanical mixing in industrial scale plants is not feasible, but some flow restrictors that avoid channelling in the extraction bed are technically possible. An extended description on modelling mass transfer may be found in Brunner (1994). Energy consumption Energy consumption for an extraction process using SC solvents may be determined from the TS diagram as described in section 1.3. 1.4.2 Applications for extraction of solids using SCF There are numerous applications for extraction of solids using supercritical fluids. Several overviews are available (Brunner, 1994; Catchpole et al., 2009; Diaz-Reinoso et al., 2007; Eltringham and Catchpole, 2007; Fang et al., 2007; Gardner, 1993; King and Srinivas, 2009; Lack and Seidlitz, 1993; Lack and Simandi, 2001; Li, 2007; Meireles, 2007; Mendes, 2007; Moyler, 1993; Mukhopadhyay, 2007; Reverchon and Marco, 2007; Stahl et al. 1987; Teelli, 2009; Temelli et al., 2007. The web pages of equipment producers (Natex, Nova Swiss, Sitec, Uhde HPT) give references to their plants. From this information, it may be seen that the highest capacity equipment is installed for coffee and tea decaffeination. The second largest application is for the extraction of hop compounds. Extraction of spices for the production of oleoresins and the extraction of bioactive compounds from plants are also very widely used applications of SC fluids. One of the latest applications is the extraction of oil from the degumming residue to obtain highly concentrated and very pure lecithin (plant designed, manufactured and erected by Uhde HPT).
1.5 Extraction of liquids using SCF There are fewer industrial units using SCF for the separation of components from liquid mixtures. Extraction from liquid mixtures by SCF is similar to liquid–liquid extraction, where compressed gas is used instead of an organic solvent. In liquid–SCF extraction processes, pressure plays an important role. When pressure and/or temperature is changed, the physicochemical properties of the SCF, such as density, viscosity, surface tension, and dielectric constant are also changed. Selective extraction of components or the fractionation of total extracts is possible by the use of different gases for the different processes and/or by changing the process parameters. Another advantage is that depending on the feed material, the density difference between the two counter-current flowing phases can be adjusted. One of the most important advantages of using supercritical fluids is the simplicity of solvent regeneration in comparison with liquid–liquid extraction,
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Principles of supercritical fluid extraction and applications 31 where in most cases, a re-extraction or distillation step is necessary; this step is energy consuming, thereby making the process more costly. Heat treatment of the extract or raffinate phase may cause the degradation of heat-sensitive substances. In extraction plants where SCFs are used, solvent regeneration is achieved by changing the pressure and/or the temperature after the extraction stage, thus changing the density and therefore the solvent power of the gas, which can be recycled after the separation of the solute. Compared with the extraction of solids with SCF, liquids may be continuously introduced and withdrawn from the high-pressure extraction unit. This offers the advantage of higher throughputs in continuous operating counter-current processes. There have been some laboratory-scale studies on extraction in systems using liquid/supercritical fluid. Data on binary systems liquid/SCF can be found, but there is less information on systems using liquid/liquid/supercritical fluid, which is necessary for the design of extraction processes of liquid mixtures with supercritical fluids. 1.5.1 Operation methods and apparatus As in conventional continuous liquid–liquid extraction in liquid/sub- or supercritical solvent, several operating modes of extraction are available. Single-stage extraction is the simplest and is used in systems where the separation factors for a solute are high. Multistage separation is necessary when the separation factor between the components is in the order of 1–10. Various modes of operation in multistage processes are used, such as multistage crossflow in which a relatively low loading of solvent with the extract is obtained in each stage. In multistage counter-current extraction, high loading of solvent with extracts is possible and a different configuration of the apparatus is possible. Counter-current liquid/sub- or supercritical fluid extraction (SFE) may be modelled by the use of typical commonly used basic equations: mass balance, energy balance, equilibrium distribution coefficients, and mass transfer rate equations. For extraction, the following data are necessary: ∑ ∑
determination of the number of theoretical stages/transfer units, size and type of a separation device in respect to separation performance, ∑ design of the solvent cycle. From the above facts and experimental data, the costs of separation using liquid/sub- or supercritical processes may be determined. The costs per tonne of the feed are influenced by the throughput and mode of operation (batch processes have higher operating costs, whereas in continuous processes the costs are lower) and are in the range from c. 607/kg feed at a throughput of c. 200 tonne per year in batch processes, down to approximately 0.067/ kg feed at a throughput of c. 60,000 tonne per year for continuous process. Brunner (2009) reported that the usual methods of calculation yield results
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32 Separation, extraction and concentration processes with an error of ±30%. Even after completion of a project, it is unlikely that the margin of error in determining costs will be less than 5%. 1.5.2 Applications The applications of liquid/sub- or supercritical fluid extraction are numerous and are used in the separation of ethanol from water (Hsu and Tan, 1994; Knez et al., 1994), the separation of aromas in various alcoholic beverages (Gamse et al., 1999), the separation of components from citrus oils (Knez, 1989) and for the purification of tocopherols (Fleck et al., 2000). The separation of caffeine from CO2 is used widely in decaffeination processes. In Fig. 1.16, the high-pressure column of an industrial-scale decaffeination process is shown. In the future, further limitations on the use of organic solvents and the demands of new applications will be the deciding factors for sustainable processing.
1.6 Conclusion SCF-based technologies offer important advantages over organic solvent technology, such as ecological friendliness and ease of product fractionation. The extraction of hop components and the decaffeination of tea and coffee are the largest scale extraction processes using sub- or supercritical solvents to be used industrially. There are also several industrial plants in operation
Fig. 1.16 Column for liquid–SCF extraction under preparation for pressure test (length 12 m, diameter 0.8 m, operation pressure 500 bar, operation temperature 100 °C. (Photo: courtesy of Uhde HPT, Hagen, Germany.)
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Principles of supercritical fluid extraction and applications 33 for the extraction of spices for the food industry and of natural substances for use in cosmetics. There are fewer industrial units engaged in the separation of components from liquid mixtures using sub- or supercritical fluids. The main advantages of using SCF for the isolation of natural products are: solvent-free products, no by-products and a low temperature in the separation process. In addition, the processes may be easily linked with direct micronisation and crystallisation from SC CO2 by fluid expansion. However, the most important advantage of the use of SCF, is the selective extraction of components or the fractionation of complete extracts. This is made possible by the use of various gases for the isolation/fractionation of components and/or by changing the process parameters. In addition to the wide use of gas for sub- or supercritical extraction, usually CO2, other sub- or supercritical solvents are also used. Sub- and supercritical CO2 and supercritical H2O are non-carcinogenic, non-toxic, non-mutagenic, nonflammable and thermodynamically stable. In addition, CO2 does not usually oxidise substrates and products, thus allowing the process to be operated at low temperatures. Water is currently the cheapest solvent and several substances are highly soluble in water. Further research on the use of subor supercritical water for the isolation and fractionation of substances is underway. One of the major process benefits is derived from the thermophysical properties of SCF: high diffusivity, low viscosity, density, and the dielectric constant of SCF, all of which may be fine tuned through changes of operating pressure and/or temperature. The limitation on further applications for obtaining extracts through highpressure technology is the price of the product; i.e. the price is relatively higher than for conventionally obtained extracts. The legal limitations on solvent residues and solvents (in products used for human applications) and the isolation/fractionation of special components from total extracts in combination with different formulations (controlled release for example) (Reverchon 2009; Weidner, 2009), chromatography (Taylor, 2009) and sterilisation processes, will lead to an increase in the use of dense gases for extraction applications. Figure 1.17 presents the number of compressed fluid extraction plants, by region (Lütge and Schuetz, 2007). It is evident that the number of extraction units will increase with time and the most ‘dense’ areas are in Europe and Asia, where a number of newly installed SCF extraction units may be observed. Figure 1.18 shows the pressures used for research and development. There is no clear trend towards higher or lower pressure. However, as part of research and development and in the filing of patents for SC extraction, both higher and lower pressures are used. Many substances have limited solubility at moderate pressure and the solubility of several substances increases by some orders of magnitude when using pressures over 500 bar. Because of this, the ultra high pressure range (up to 2500 bar) enables additional fractionation of substances in the extract
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34 Separation, extraction and concentration processes 50 Europe
America
Asia
Australia
45 40 35
Number
30 25 20 15 10 5 0 1980
1985
1990
Year
1995
2000
2005
Fig. 1.17 Number of SFE plants having a total extraction volume >500 L by region. (Source: Schuetz Consulting/Uhde HPT.) 800 700
Pressure (bar)
600 500 400 300 200 100 0 1980
1985
1990
1995 Year
2000
2005
2010
Fig. 1.18 Maximum pressure of SFE processes. (Source: Schuetz Consulting/Uhde HPT.)
by pressure- and/or temperature-dependent precipitation. In Fig. 1.19 an ultra high pressure extraction unit operating up to 2500 bar is presented. The evaluation of costs for several plant materials processed by ultra high pressure extraction using SC CO2 is shown in Fig. 1.20. It is clear that with © Woodhead Publishing Limited, 2010
Principles of supercritical fluid extraction and applications 35
E2
S1
E1
S2
Fig. 1.19 Ultra high pressure extraction unit (operating pressure 2500 bar). (E1 and E2, extractors; S1 and S2, separators). (Photo: courtesy of Uhde HPT, Hagen, Germany.) 2.50
Specific processing costs (7/kg)
Operating costs (7/kg) Investment (7/kg) 2.00
1.50
1.00
0.50
0.00
300
500 1000 Pressure (bar)
1500
Fig. 1.20 Specific processing costs versus operating pressure (Lütge et al., 2007)
an increased operating pressure in a SC extraction plant (retaining the same capacity of equipment), the total processing costs decrease. In the future, further limitations on the use of organic solvents, new applications of several substances, changing customer requirements, sustainable © Woodhead Publishing Limited, 2010
36 Separation, extraction and concentration processes production and processing of substances, will all lead to new developments in high-pressure processing. We can be sure that advances in the field of high-pressure research into cheap and environmental friendly solvents, such as CO2 and some other gases and sub- or supercritical water, will open up new pathways for substances and products produced at high pressure.
1.7 References Y. Adachi and H. Sugie, Fluid Phase Equilib. 28(2) (1986) 103–118. A. Anderko, Fluid Phase Equilib. 61 (1990) 145–225. A. Anderko, in: J.V. Sengers, R.F. Kayser, C.J. Peters, H.J. White (Eds.), Equations of state for fluids and fluid mixtures, Elsevier, Amsterdam, 2000, pp. 75–126. J.D.S. Arons and G.A.M. Diepen, Rec. Trav. Chim. Pays-Bas, 82 (1963) 249–256. A. Bertuco and G. Vetter, High pressure technology: fundamentals and application, Industrial Chemistry Library, volume 9, 2001. G. Brunner, Counter-current separations, J. Supercrit. Fluids 47 (2009) 574–582. G. Brunner, Gas extraction. An introduction to fundamentals of supercritical fluids and the application to separation processes, Darmstadt: Steinkopff; New York: Springer, 1994. O.J. Catchpole, S.J. Tallon, W.E. Eltringham, J.B. Grey, K.A. Fenton, E.M. Vagi, M.V. Vyssotski, A.N. MacKenzie, J. Ryan, Y. Zhu, The extraction and fractionation of specialty lipids using near critical fluids, J. Supercrit. Fluids 47 (2009) 591–597. R.P. Danner and P.A. Gupte, Fluid Phase Equilib. 29 (1986) 415–430. B. Diaz-Reinoso, A. Moure, H. Dominguez, and J.C. Parajó, Antioxidant extraction by supercritical fluids, in J.L. Martinez, Supercritical fluid extraction of nutraceuticals and bioactive compounds, CRC Press, 2007, 275–304. R. Dohrn, Berechnung von Phasengleichgewichten, Friedr. Vieweg & Sohn Verlagsgesellschaft mbH, Braunschweig, 1994. R. Eggers, Supercritical fluid technology in oil and lipid chemistry, AOCS Press, 1996, 35–62. W. Eltringham and O.J. Catchpole, Processing of fish oils by supercritical fluids, in J.L. Martinez, Supercritical fluid extraction of nutraceuticals and bioactive compounds, CRC Press, 2007, 141–188. T. Fang, M. Goto, M. Sasaki, and D. Yang, Extraction and purification of natural tocopherols by supercritical CO2, in J.L. Martinez, Supercritical fluid extraction of nutraceuticals and bioactive compounds, CRC Press, 2007, 103–140. U. Fleck, G. Brunner and R. Karge, Purification of synthetic crude tocophenol acetate by means of supercritical fluid extraction, in: Proceedings of the 5th International Symposium on Supercritical Fluids, Atlanta, April 2000. T. Gamse, I. Rogler and R. Marr, 1999, Supercritical CO2 extraction for utilisation of excess wine of poor quality, J. Supercrit. Fluids 14 (1999) 123–128. T. Gamse and R. Marr in A. Bertucco, G. Vetter. High pressure process technology: fundamentals and applications, Industrial Chemistry Library, volume 9, 2001, 383. D.S. Gardner, Commercial scale extraction of alpha acids and hop oils with compressed CO2, in M.B. King and T.R. Bott, Extraction of natural products using near-critical solvents, Chapman & Hall, 1993, 84–100. J.-H. Hsu and C.-S. Tan, Separation of ethanol/water solution with supercritical CO2 in the presence of a membrane, in S.S.H. Rizvi, Supercritical fluid processing of food and biomaterials, Chapman & Hall, 1994, 114–122. M.J. Huron and J. Vidal, Fluid Phase Equilib. 3 (1979) 255.
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Principles of supercritical fluid extraction and applications 37 P.G. Jessop and W. Leitner, Chemical synthesis using supercritical fluids, Wiley-VCH, July 1999. J.W. King and K. Srinivas, Multiple unit processing using sub- and supercritical fluids, J. Supercrit. Fluids 47 (2009) 598–610. Ž. Knez, Separation of components of citrus oils – industrial project, unpublished data, 1989. Ž. Knez, F. Posel and I. Krmelj, High pressure extraction of organics from water, in S.S.H. Rizvi, Supercritical fluid processing of food and biomaterials, Chapman & Hall, 1994, 181–186. Ž. Knez and M. Škerget, Phase equilibria of the vitamins D2, D3 and K3 in binary systems with CO2 and propane, J. Supercrit. Fluids 20 (2001) 131–144. P.H. van Konynenburg and R.L. Scott, Critical lines and phase equilibria in binary van der Waals mixtures, Phil. Trans. Roy. Soc. London, Ser. A 298 (1980) 495. E. Lack and H. Seidlitz, Commercial scale decaffeination of coffee and tea using supercritical CO2, in M.B. King and T.R. Bott, Extraction of natural products using near-critical solvents, Chapman & Hall, 1993, 101–139. E. Lack and B. Simandy in A. Bertucco and G. Vetter. High pressure technology: fundamentals and application, Industrial Chemistry Library, volume 9, 2001, 537–575. S. Li, Application of supercritical fluids in traditional chinese medicines and natural products, in J.L. Martinez, Supercritical fluid extraction of nutraceuticals and bioactive compounds, CRC Press, 2007, 215–242. T.W. de Loos, On the phase behaviour of asymmetric systems: The three-phase curve solid–liquid–gas. J. Supercrit. Fluids 39 (2006) 154–159. D. Luedecke and J.M. Prausnitz, Fluid Phase Equilib. 22(1) (1985) 1–19. C. Lütge and E. Schuetz, Market trends and technical developments in high pressure technology, in: Ž. Knez and M.J. Cocero, Proceedings of the 5th international symposium on high pressure process technology and chemical engineering, June 24–27, 2007, Segovia, Spain. C. Lütge, M. Bork, Ž. Knez, M. Knez Hrn�i�, M. Krainer, Ultra high pressure dense gas extraction and fractionation, in: Ž. Knez and M.J. Cocero, 5th International symposium on high pressure process technology and chemical engineering, June 24–27, 2007, Segovia, Spain. R. Marr and T. Gamse, High pressure technology: fundamentals and application, Industrial Chemistry Library, volume 9, 2000, 396–402. M.A. McHugh and V.J. Krukonis, Supercritical fluid extraction: principles and practice, Butterworths, Stoneham, 1986. M.A. Meireles, Extraction of bioactive compounds from Latin American Plants, in J.L. Martinez, Supercritical fluid extraction of nutraceuticals and bioactive compounds, CRC Press, 2007, 243–274. R. Mendes, Supercritical fluid extraction of active compounds from algae, in J.L. Martinez, Supercritical fluid extraction of nutraceuticals and bioactive compounds, CRC Press, 2007, 189–214. D.A. Moyler, Extraction of flavours and fragrances with compressed CO2, in M.B. King and T.R. Bott, Extraction of natural products using near-critical solvents, Chapman & Hall, 1993, 140–183. M. Mukhopadhyay, Processing of spices using supercritical fluids, in J.L. Martinez, Supercritical fluid extraction of nutraceuticals and bioactive compounds, CRC Press, 2007, 337–366. J.P. O’Connell, W. Weber and G. Brunner, Measurement and thermodynamics of triglyceride melting in near-critical fluids, in: Paper Presented at the AIChE Annual Meeting, Indianapolis, USA, 2003. A.Z. Panagiotopoulos, R.C. Reid, S. Watanasiri, Fluid Phase Equilib. 29 (1986) 525.
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38 Separation, extraction and concentration processes J.M. Prausnitz, R.N. Lichtenthaler and E. Gomes de Azevedo, Molecular thermodynamics of fluid phase equilibria. Prentice-Hall, Inc., Englewood Cliffs, New Jersey, 1986. E. Reverchon and I. De Marco, Essential oils extraction and fractionation using supercritical fluids, in J.L. Martinez, Supercritical fluid extraction of nutraceuticals and bioactive compounds, CRC Press, 2007, 305–336. E. Reverchon, R. Adami, S. Cardea, G. D. Porta, Supercritical fluids processing of polymers for pharmaceutical and medical applications, J. Supercrit. Fluids 47 (2009) 484–492. R.J. Sadus, High pressure phase behavior of multicomponent fluid mixtures, Elsevier Amsterdam, 1992. S.I. Sandler and H. Orbey, Mixing and combining rules; in: J.V. Sengers, R.F. Kayser, C.J. Peters, H.J. White (Eds), Equations of state for fluids and fluid mixtures, Elsevier, Amsterdam, 2000, pp. 321–357. R. Sandoval, G. Wilczek-Vera, J.H. Vera, Fluid Phase Equilib. 52 (1989) 119. J. Schwartzentruber and H. Renon, Ind. Eng. Chem. Res. 28 (1989a) 1049. J. Schwartzentruber, H. Renon, Fluid Phase Equilib. 52 (1989b) 127. H. Sovova, Mathematical model for supercritical fluid extraction of natural products and extraction curve evaluation, J. Supercrit. Fluids 33 (2005) 35–52. E. Stahl, K.-W. Quirin, D. Gerard, Verdichtete Gase zur Extraktion und Raffination, Springer-Verlag, 1987, 82–225. M Škerget, Z. Novak-Pintari�, Ž Knez and Z. Kravanja, Estimation of solid solubilities in supercritical carbon dioxide, Peng–Robinson adjustable binary parameters in the near critical region. Fluid Phase Equilib. 5086 (2002) 1–22. L.T. Taylor, Supercritical fluid chromatography for the 21st century, J. Supercrit. Fluids 47 (2009) 566–573. F. Temelli, M.D.A Saldaña, P.H.L. Moquin, and M. Sun, Supercritical fluid extraction of specialty oils, in J.L. Martinez, Supercritical fluid extraction of nutraceuticals and bioactive compounds, CRC Press, 2007, 51–102. F. Temelli, Perspectives on supercritical fluid processing of fats and oils, J. Supercrit. Fluids 47 (2009) 583–590. W.H. Tuminello, G.T. Dee and M.A. McHugh, Macromolecules 28 (1995) 1506. E. Weidner, V. Wiesmet, Z. Knez, M. Skerget, Phase equilibrium (solid–liquid–gas) in polyethylene glycol–carbon dioxide systems, J. Supercrit. Fluids 10 (1997) 139–147. E. Weidner, High pressure micronization for food applications, J. Supercrit. Fluids 47 (2009) 556–565. D.S.H. Wong and S.I. Sandler, AIChE J. 38 (1992) 671. K. Zosel, Process for the decaffeination of coffee, US Patent 4247570, January, 1981.
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Principles of pressurized fluid extraction and applications 39
2 Principles of pressurized fluid extraction and environmental, food and agricultural applications C. Turner and M. Waldebäck, Uppsala University, Sweden
Abstract: The main principles of pressurized fluid extraction (PFE), including basic extraction theory and the effects of solvent selection and temperature variations are discussed. Methods for achieving selectivity during the extraction are described, and future trends explored. Examples of applications include extraction methods developed for environmental and food analysis. Some advice on building equipment in the laboratory is given. Key words: food analysis, pressurized fluid extraction, PFE, environmental analysis, solvent extraction.
2.1 Introduction When developing new products, processes and technologies, it is important to strive for sustainable development. This is especially important because chemists, chemicals, and the chemical industry as a whole are commonly regarded to be the cause of many of the current environmental problems. To ensure sustainability, new processes and new techniques should be studied from a life cycle point of view. Life cycle assessment (LCA) is a methodology in which the entire life cycle of a product or utility effect is analyzed, such as extraction and processing of raw materials, production, distribution, use, consumption and disposal as well as the potential ecological effects. In addition, energy conversions occurring in a life cycle and the resulting burden on the environment are assessed. The indoor environment and health perspective is as important as the outdoor environment. Chemicals should be used in smaller amounts and the ones used should be less hazardous.
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40 Separation, extraction and concentration processes One process in which large volumes of organic solvents are used is the procedure of extraction. Traditional liquid–liquid extraction (LLE), also known as solvent extraction, and liquid–solid extraction (LSE) are techniques in which different compounds can be separated from each other based on their relative solubility. For a solid sample, the separation of a substance from the mixture occurs by dissolving that substance in an appropriate solvent. The extraction process usually requires several hours or even several days to perform, depending on the extraction temperature. These processes are slowly being replaced by more attractive alternatives. The most widely used extraction techniques today are still Soxhlet (developed in 1879) and sonication extraction (from 1960). These classical techniques are usually multi-step procedures based on exhaustive extractions from a sample matrix followed by successive clean-up steps before analysis. Such sample preparation procedures require large amounts of sample, sorbents and organic solvents, which are often hazardous and/or toxic, resulting in high costs of both purchase and disposal (Ramos et al., 2002). These methods also demand extensive manual handling, which often creates work-related health problems. During the past few decades new techniques have been developed and among the more successful ones are those employing high-diffusion liquids. Highdiffusion liquids are in this chapter defined as liquids of elevated temperature and pressure. By using high-diffusion liquids, the diffusion coefficient of the liquid is increased. This is the most effective way to increase the rate of the extraction process and decrease the required amount of organic solvent. Diffusion rates in liquids have been shown to increase about 2–20 fold upon increasing the temperature from 25 to 150 °C (Perry et al., 1984). Thus, the mass transfer rate can be increased and the whole extraction process becomes faster. Techniques that employ high-diffusion liquids with raised temperature and pressure are microwave-assisted extraction (MAE) and pressurized fluid extraction (PFE) (Majors, 1996). Another related extraction technique is supercritical fluid extraction (SFE), which usually employs carbon dioxide at pressures and temperatures above 74 atm and 31 °C, respectively, resulting in a liquid-like density in combination with high diffusion rates and low viscosity. In PFE, the solvent is kept in a liquid phase even at temperatures much above the atmospheric boiling point as a result of the applied pressure. This technique is also known as pressurized liquid extraction (PLE“), pressurized solvent extraction (PSE™), accelerated solvent extraction (ASE“) and enhanced solvent extraction (ESE). PFE was originally introduced at the Pittsburgh conference (Pittcon) in 1995 as ASE“ by the Dionex Corporation. The PFE technique utilizes the same basic principles as traditional liquid solvent extractions, but the extractions are carried out at higher temperatures and pressures. An increased temperature during the extraction gives more efficient extraction, and results in both time savings and lowered solvent consumption (Richter et al., 1996). The principle of PFE is similar to
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Principles of pressurized fluid extraction and applications 41 MAE, where the solvent is heated by microwave energy, although in the PFE technique both higher temperatures and pressures can be obtained, independently. In this chapter, the basic instrumentation of dynamic (continuous flow) and static (batch mode) PFE as well as its principles of operation are described. Furthermore, fundamental extraction theory is explained including illustrative examples and classical extraction models from literature. A guide to solvent selection is included, which on an elementary level guides the reader to an appropriate choice of solvent in various applications, depending on the chemistry of the sample matrix and analytes to be extracted. Effects of changing the extraction temperature, time and pressure in PFE are described, in addition to some remarks on how to obtain selectivity as well as accuracy and precision. One section is devoted to describing common applications for this technique. Finally, future trends for PFE in terms of sustainability, ‘like-nature’ applications, hyphenated techniques, downscaling and transfer to industry are depicted. The final paragraph also lists commercial vendors of PFE equipment and some advice for building equipment in the laboratory. The next subsection is largely based on a doctoral thesis written by one of the authors of this chapter (Waldeback, 2005).
2.2 Instrumentation and principles of pressurized fluid extraction 2.2.1 Basic instrumentation Extractions using a pressurized liquid above its atmospheric boiling point require open/close valves and/or pressure restrictors, in order to maintain pressure during the extraction. Two main types of instrumentation can be used: static PFE, which is a batch process with one or several extraction cycles with replacement of solvent in between; and dynamic PFE, in which the extraction solvent is continuously pumped through the extraction vessel containing the sample. A schematic of a static PFE instrument is shown in Fig. 2.1. Temperatures applied usually range from room temperature to 200 °C and pressures are generally between 35 and 200 bar. Filter paper is inserted into a stainlesssteel extraction cell followed by the sample, if necessary mixed with a drying agent. The cell is either loaded on a carousel and automatically placed in the oven, or, for simpler equipment, manually placed into the oven. Most equipment has a pump that fills the extraction cell with solvent. Commonly, a static valve in combination with a pressure relief valve control the pressure in the sample vessel during the static extraction; either by adding more solvent to the cell or by opening the static valve, whichever is needed to maintain the desired pressure. The first part of the extraction is a pre-heating step to reach thermal equilibrium. During this heating, thermal expansion of the
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42 Separation, extraction and concentration processes Purge valve
Pressure relief valve Extraction cell Oven
Pump
Solvent A Solvent B Static valve
Vent Solvent C Nitrogen
Waste vial
Solvent D Collection vial
Fig. 2.1 Schematic diagram of a PFE system (courtesy of Dionex Corporation).
solvent occurs and causes an increase in pressure within the cell. When the set values are achieved static extraction is performed during a selected time, typically 5–10 min. After the static time, part of the solvent in the extraction cell is replaced with fresh solvent, to start the next extraction cycle. After the last cycle, the sample cell is purged with an inert gas such as nitrogen to remove the remaining solvent from the cell and the lines to a vial that contains the extract. In some equipment, sequential extractions can easily be performed by repeating the procedure with a new solvent, and purging in the same or a different collection vial. Dynamic PFE is quite similar to static PFE, but requires a more sophisticated high-pressure or HPLC pump as well as a pressure restrictor rather than a static open/close valve. The equipment is similar to that used for ‘superheated water chromatography’ (Smith et al., 1999), but the column is replaced with an extraction cell and the tubing has slightly wider inner-diameter. Currently, there is no dynamic PFE equipment available on the market. In fact, there are only a few different types of commercial instruments available (Section 2.5.1), but many examples of home-made instrumentation have been described. Some of them perform both dynamic and static extractions at temperatures above 200 °C (Bautz et al., 1998; Hawthorne et al., 2000; Lou et al., 1997; Vandenburg et al., 1998). A short paragraph about building
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Principles of pressurized fluid extraction and applications 43 your own equipment can be found in Section 2.5.2, including Fig. 2.8 that shows a schematic of what a dynamic PFE system may look like. 2.2.2 Extraction strategy In order to perform efficient and smart extractions, an understanding of the matrix characteristics and the different steps of the extraction is of great importance. The nature of the sample matrix (water and organic/inorganic content) and its physical characteristics (homogeneity, porosity, particle size) should also be considered. To better understand the extraction process, two models showing the distribution of analytes in different types of sample matrices are illustrated below. Figure 2.2 is a conceptualization of an aggregate of matrix particles from a source-separated household waste, and the possible sites where analytes in this instance chlorinated paraffins, are expected to be found (Nilsson et al., 2001). Figure 2.2 works as a model showing the variety of possible (and likely) positions and status of analytes in many different types of sample matrices, i.e. the analyte can be: 1. 2. 3. 4. 5.
adsorbed at the surface of the matrix, dissolved in the pore solvent and/or adsorbed at the pore surface, dissolved/adsorbed in a micro/nano pore, chemically bonded to the matrix, or dissolved in the bulk solution.
A simpler extraction model can be described for extracting small organic compounds from polymer particles, for instance the antioxidant Irganox
5 1 2
4 3
Household waste particle Stagnant solvent layer
Fig. 2.2 Schematic of a household waste particle and some possible sites (for explanation of numbers see text) where the analyte (chlorinated paraffin) might be adsorbed or chemically bonded (Waldeback, 2005).
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44 Separation, extraction and concentration processes 1076 from the polymer LLDPE (Waldeback et al., 1998), Fig. 2.3. The extraction process of compounds from polymer particles, or other similar sample matrices, generally follows the following steps: 1. diffusion of the solvent into the matrix, 2a. desorption of the analytes from the matrix (including breaking of chemical bonds), 2b. solvation of the analyte into the extraction solvent, 3. diffusion of the analyte out from the matrix, and 4. diffusion of the analyte through the stagnant solvent layer and into the bulk solvent. Knowledge about distribution coefficients and distribution ratios are useful tools to provide guidance to the selection of solvent for the extractive separation process. Solubility of the analytes, their diffusivity in the solvent and matrix characteristics are the main factors to consider, when choosing a solvent for a successful extraction process. It is also important to understand the mass transfer mechanism across chemical/physical interfaces in order to design liquid/liquid and liquid/solid extraction processes. Toxicity and sustainability aspects of the solvent should also be considered. Additionally, in most of the analytical-scale applications, the concentration of the target molecule is very low, and thus the rate of the extraction is not limited by the analyte concentration in the extraction solvent, but rather determined by the rate of mass transfer out of the matrix. In order to perform a fast and complete extraction, a solvent has to be chosen that has the right chemical properties to dissolve and release the analyte, but should preferably not dissolve other solutes in the sample, i.e. the solvent power should not be higher than needed. Solubility theory has been discussed and proposed in classic works by J. Hildebrand, who combined the correlation between vaporization and intermolecular forces, van der Waal forces and hydrogen bonding, to the correlation between vaporization and solubility behaviour (Hildebrand and Scott, 1962; 1964). This model assumes that the same intermolecular attractive forces have to be overcome to vaporize a liquid as to dissolve an analyte. The term ‘solubility parameter’ d was described by Hildebrand as the square root of the cohesive energy density c, giving a numerical value indicating the analyte behaviour in a specific solvent, equation [2.1], where DH is the heat of vaporization (J mol–1) of the solvent, R the gas constant (J K–1 mol–1), T the temperature (K) and V is the molar volume of the analyte.
d = c = DH – RT [2.1] V Hansen (1967) took Hildebrand’s work further and assumed that the total cohesive energy is a linear addition of three components; dh (hydrogen bonding ability contribution), dd (dispersion coefficient contribution), and
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Principles of pressurized fluid extraction and applications 45 dp (polarity contribution). They are linked by equation [2.2], where dt is the total solubility parameter (Fitzpatrick and Dean, 2002).
d t2 = d h2 + d d2 + d p2
[2.2]
Fitzpatrick and Dean (2002) predicted a suitable solvent to extract persistent organic pollutants (POPs) from contaminated soil and certified reference material using the Hildebrand solubility parameter and confirmed the results by experiments. The ideal extraction solvent from the calculations was a mixture of acetonitrile and dichloromethane (1:1 v/v), perhaps the best solvent from the view of solvent power, but not from a health and environmental view. When choosing a solvent it is a good start to have the Hildebrand and Hansen’s theories in mind. However, the choice of solvent in a particular situation involves other factors apart from the solvent power. As described in Fig. 2.2 and 2.3, the solvent has to penetrate the matrix thoroughly, break the bonds between the matrix and the analytes of interest, help the dissolved analytes to diffuse out from the matrix and finally be dissolved in the extraction solvent. The process of solute transfer across an interface between two liquid phases may be rate-controlled by molecular diffusion, by motion of eddies, by irregular surface disturbances or even by chemical reactions in the bulk of a phase or in the interface region. Local velocities in the interface region could be of importance as well as other factors affecting the local conditions, not least the presence of surface-active agents. The dependence of matrix geometry has been reported for SFE of mineral oils applied to metal devices (Bjorklund et al., 1996). Irganox 1076
1
2 3
Polymer particle Stagnant solvent layer
4
Fig. 2.3 Schematic showing the extraction of Irganox 1076 from an LLDPE polymer particle (Waldeback, 2005).
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46 Separation, extraction and concentration processes Partitioning processes have a central role of concern in the extraction procedure. These involve partitioning of the analytes between the surface of the matrix and the solvent, as well as chemisorption of the analytes on active surface sites and within the solvent. Different matrices behave somewhat differently, e.g. polymeric samples usually build up a layer of stagnant liquid around the polymeric particles, as seen in Fig. 2.3, through which the analyte has to transfer into the extracting solvent. In this case, the partitioning of the analyte between the stagnant liquid and the extraction solvent has to be considered. Soils differ strongly in surface physicochemical properties and grain-size characteristics. Sediments, on the other hand, contain water having various types of bonding, from free water available for the plants, to water strongly bound to the particles, therefore a variety of equilibria take place. In milk powder, hard shells of lactose are formed during the short crystallization time when the milk is spraydried (Walstra and Jenness, 1984), and this may lead to difficult extraction of lipids and fat soluble vitamins. If the extraction solvent is nonpolar, it may be necessary to add a more polar alcohol such as methanol or ethanol to the sample before extraction (Turner and Mathiasson, 2000). There are only a few theoretical models suggested for the PFE technique. For instance, Vandenburg et al. applied the ‘hot ball’ model, originally described for SFE, for the extraction of additives from polymeric samples using PFE (Vandenburg et al., 1998). By plotting ln(m1/m0), where m1 is the mass of analyte remaining in the particle of radius r at time t, m0 is the initial amount of analyte and D the diffusion coefficient of the analyte in the solvent, a linear portion is given in equation [2.3] (Cotton et al., 1993). In this instance, when ln(m1/m0) is plotted against time, the line falls steeply initially and shortly after becomes linear where it follows equation [2.3] (Vandenburg et al., 1998):
ln(m1/m0) = – 0.4977 – (p2Dt/r2)
[2.3]
The physical explanation of the shape of the curve is that the analyte near the surface is rapidly extracted until a smooth falling concentration gradient is established across the particle. The extraction rate is then completely controlled by the rate at which the analyte diffuses to the surface. By plotting the amount of extracted analyte versus the extraction time for different solvents at different temperatures, the resulting curves showed a good fit to the ‘hot ball’ model. With a good kinetic model of the extraction process it would be possible to predict experimental parameters, and find out where the extraction is expected to be only diffusion dependent. However, the models mentioned here can at most be considered to give useful hints when developing new extraction methods based on PFE. There is no theoretical model that includes strong solute-matrix interactions caused by chemical bonding between the solute molecules and active sites on the matrix. In addition, sample matrices are seldom homogeneous, thus the penetration of the solvent is difficult to foresee. As a result, accurate descriptions are by
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Principles of pressurized fluid extraction and applications 47 no means easy to provide because of experimental difficulties, as well as the chemical and mathematical complexity of the total extraction process. Empirical approaches, in combination with multivariate chemometric methods, are the most widely used. 2.2.3 Solvent selection As mentioned in Section 2.2.2, a solvent has to have the right chemical properties to desorb and dissolve the analyte preferably without dissolving other solutes in the sample, i.e. the solvent power should not be higher than needed. Richter et al. (1996) have suggested the conventional LLE solvent to be used when developing a new PFE method. Generally, when choosing a solvent the rule-of-thumb is ‘like dissolves like’, i.e. polar solvents dissolve polar analytes, and nonpolar solvents dissolve nonpolar analytes. In addition, the dipole moment and/or dielectric constants of solvents are useful for selecting an appropriate extraction solvent, see Table 2.1. Other important aspects to consider when facing a new extraction problem/challenge is in how the analyte of interest is dissolved or attached Table 2.1 Chemical properties of common extraction solvents (from: Handbook of chemistry and physics, CRC press, 52nd Edition, 1971–1972; and Handbook of chemistry and physics, CRC Press, 60th Edition, 1979–1980) Solvent Boiling Density point (°C) (g mL–1)
Vapour Dipole pressure moment at 20 °C (debye) (mbar)
Dielectric constant at 20 °C (mbar)
Polar, protic solvents Water 100.0 Methanol 64.9 Ethanol 78.5 n-Propanol 97.1 2-Propanol 82.4 n-Butanol 117.7
23 128 44 – 43 –
1.85 1.70 1.69 1.68 1.66 1.66
80.2 32.2 24.3 20.1 18.3 17.8
3.96
47.2
1.000 0.791 0.789 0.803 0.785 0.810
Polar, aprotic solvents Dimethyl sulfoxide 189 (DMSO) Acetonitrile 81.6 Acetone 56.2 Dichloromethane 40.0 Tetrahydrofuran 66 Ethyl acetate 77.1
0.786 0.790 1.326 0.886 0.894
93 233 453 200 97
3.92 2.88 1.60 1.63 1.78
37.5 20.7 9.1 7.4 6.0
Nonpolar solvents Chloroform Diethyl ether Toluene Benzene n-Hexane
1.498 0.714 0.867 0.879 0.660
211 507 29 108 160
1.0 1.15 0.4 0 0
4.8 4.3 2.4 2.3 1.9
61.7 34.5 110.6 80.1 68.9
1.092
0.56
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48 Separation, extraction and concentration processes to the matrix, see Fig. 2.2 and 2.3, and also the type and strength of the intermolecular forces between the solvent molecules, which partly can be elucidated by the ΔHvap value, alternatively by the vapour pressure of the solvent (Table 2.1). When dissolving a solute, the intermolecular forces, i.e. hydrogen bonding, dipole–dipole and/or van der Waals interactions between the solvent molecules have to break and form new bindings/interactions with the solute. In instances when the solute is chemically bonded to the matrix, as in Fig. 2.2 (site 4), the extraction solvent has to overcome such bindings and then form new solute–solvent interactions, which altogether will lead to a more stable system of lower energy. Methanol and acetonitrile, two organic solvents commonly used in extraction and separation processes, have somewhat similar dielectric constants, 32 and 37, respectively, but differ strongly with regard to intermolecular forces as methanol forms hydrogen bonds, whereas acetonitrile relies on dipole–dipole interactions. Hence, acetonitrile, which is a less environmentally friendly solvent than methanol, should only be used in cases where such dipole–dipole interactions are more suitable. What are the partitioning processes or equilibrium reactions taking place during the extraction? For example, when extracting samples of high water content, such as wet sediments, vegetables or animal muscles, using organic solvents, there is always a risk that a nonpolar solvent will not be able to penetrate water-sealed pores that contain the analyte, and thus result in a lower extraction yield and unrepeatable results. The more knowledge about the chemistry of the target solutes, pertinent coextractable (unwanted) compounds and chemical and physical properties of the sample matrix and solute–matrix interactions, the easier it is to make a good selection of extraction solvent for instance from the ones listed in Table 2.1. One interesting aspect with PFE is the possibility of choosing a solvent, apart from very low or high pH solvents that may not be applicable in all kinds of equipment owing to possible corrosion of the tubings. In a collaborative study with an oil company, the objective was to predict the amount of process chemicals that could migrate into the sea if there was an oil discharge. By using a water solution of the same salinity as seawater, oil samples containing added process chemicals were extracted at different temperatures and times. The results could be used to predict what would probably happen in a real world situation (data not presented – only used internally at the company). This type of investigation is of course more environmentally relevant than a study performed with traditional extraction techniques based on organic solvents. How does the PFE technique and the solvent selection fit to a more sustainable chemistry? One of the advantages with PFE is that owing to the higher temperature and higher mass transfer during the extraction, more environmentally friendly solvents or solvent mixtures can be used according to Nilsson et al., 2001). As described above, the solvation power of a liquid increases with higher temperature and higher pressure. Furthermore, the
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Principles of pressurized fluid extraction and applications 49 dielectric constant decreases with increasing temperature, which implies that the solvent polarity can be tuned by changing the temperature. This is especially true for water, the dielectric constant of which decreases from e = 78 at 25 °C to e = 56 at 100 °C; e = 36 at 200 °C; and e = 8 at 400 °C, since hydrogen bonding between the water molecules become less and less substantial at higher temperatures (Hawthorne et al., 1994). This can be compared to other polar, protic solvents and nonpolar solvents in Table 2.1. Hence, water is a very interesting alternative as an extraction solvent, and can potentially replace many of the organic solvents conventionally used in solvent extraction. The first work describing the use of water at elevated temperature and pressure for extraction of nonpolar solutes is one by Hawthorne et al. (1994). In this work, polar organics (e.g., chlorinated phenols), low-polarity organics (e.g., PAH), and nonpolar organics (alkanes) were extracted from soils and sediments using water at temperatures ranging from 50 to 400 °C. In our own work, water at 120 °C and 50 bar was used as a solvent for the extraction of polyphenolic glycosides from onions, as compared with using a conventional water/methanol (1:1) mixture at 80 °C (Turner et al., 2006). The developed method was also faster, 15 min compared with 2 h for the conventional. In a continuous study, a life cycle assessment was conducted to assess environmental impacts of the two methodologies, showing that the production and use of methanol compared with water has a large impact in terms of carbon dioxide emissions and energy usage, whereas the difference in energy usage for heating the solvents in the two extraction methods was much smaller (Lindahl et al., 2010). If more nonpolar solutes are to be extracted, and these solutes are slightly thermolabile, then ethanol is an interesting alternative to water since not as high a temperature is needed to obtain the same solvent strength. For instance, betulin was extracted from birch bark using ethanol at 120 °C and 50 bar in only 10 min (Co et al., 2009). Initially, the study tested to see if water at elevated temperature close to 200 °C could be an appropriate solvent for betulin, because pressurized liquid water at this temperature has a dielectric constant of around 36, and this is similar to that of methanol at room temperature (e = 32, see Table 2.1), which is indeed a good solvent for betulin. However, it turned out that it is not sufficient to use the dielectric constant as a measure of solubility of a solute in a solvent, as other chemical properties are also important. For this reason, the total (sometimes called ‘relative’) solubility parameter taking dh, dd and dp into account (equation [2.2], Hansen, 2007) was plotted as a function of temperature for betulin, water and ethanol, showing that even at 250 °C, the solubility of betulin in water is estimated to be very low (Fig. 2.4), whereas ethanol at temperatures between 50 and 150 °C should be appropriate, which was confirmed by experiment (Co et al., 2009).
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50 Separation, extraction and concentration processes 50
Solubility parameter (MPa0.5)
45 40 35 30 25 20 15 Water
10
Betulin
5
Ethanol
0 0
50
100 150 Temperature (°C)
200
250
Fig. 2.4 Solubility parameters as a function of temperature for betulin, ethanol and water (Co et al., 2009).
2.2.4 Temperature effects An increase in temperature usually tends to promote solubility, as the thermal kinetic energy rises. Increasing the temperature also facilitates analyte diffusion and/or reduces interactions between analytes and the matrix by disrupting intermolecular forces such as van der Waal’s forces, hydrogen bonding and dipole attractions. Higher temperatures also decrease the viscosity of a liquid solvent, thus enabling better penetration of matrix particles. An increased temperature will also decrease the surface tension of the solvent, allowing the solvent to better ‘wet’ the sample matrix (Richter et al., 1996). Both lower viscosity and lower surface tension facilitate better contact of the solvent with the solutes and thereby enhance the extraction. Besides the selection of an extraction solvent, temperature can be considered as the second most important parameter in PFE. The advantages of using higher temperatures are described in every study published on PFE. As a general rule of thumb, a higher temperature gives a better extraction yield. It needs to be remembered that the range of solvents applicable in PFE is much wider than that of Soxhlet extraction. A poor solvent for Soxhlet extraction can, however, be a good solvent in PFE owing to the higher extraction temperatures used (Lou et al., 1997). Hence, as described below, by using a higher temperature, a more environmentally friendly and/or less toxic solvent can be used, for instance water or ethanol. However, other factors might hinder the use of the highest temperature even if it should give the best yield, e.g. as predicted in an experimental design. For example, Waldeback et al. (1998) found that the sample matrix (granules of LLDPE) started to dissolve/melt and block the tubings of the instrument depending on solvent and temperature, e.g. at temperatures above 75 °C © Woodhead Publishing Limited, 2010
Principles of pressurized fluid extraction and applications 51 using tetrahydrofuran as a solvent. Pihlstrom et al. (2002) found that, where pesticides from canola seed were extracted, the results from the screening study showed that the interpretation of the chromatograms was difficult owing to the large number of co-extracted compounds. It was obvious that higher temperatures gave more matrix peaks (Nemoto and Lehotay, 1998), which were most probably derived from co-extracted lipids. In another study (Waldeback et al., 2004), where squalene was extracted from olive biomass, the thermostability of the analyte was a limiting factor in the choice of extraction temperature. To determine the concentration of squalene in the olive biomass, an optimization study using experimental design was performed for the variables extraction temperature, extraction time and concentration of acetone in a mixture of acetone and 2-propanol. Significant factors were determined to be temperature, extraction time, the interaction between temperature and extraction time, and the square of the extraction time. This study showed that, at temperatures above 100 °C, the yield of squalene decreased with extraction times longer than 12–15 min, indicating that squalene either decomposed or reacted with other sample components or with the solvent. An example of extremely thermolabile compounds are anthocyanins, which are antioxidants found in vegetables and fruits, such as grapes, red cabbage, red onions, blueberries and red beets. PFE has been used to extract anthocyanins from red cabbage and red onions using water/ethanol/formic acid (94:5:1, vol%) at 99 °C and 50 bar (Arapitsas and Turner, 2008; Petersson et al., 2008, 2010). However, preliminary results show that anthocyanins extracted from red onion start to degrade after only a few minutes at a set temperature. Hence, extraction and degradation naturally occurs simultaneously in PFE. The kinetics of extraction and degradation depend on the nature of the sample matrix, on the chemical properties of the solutes to be extracted, and on temperature. Theoretical extraction curves were calculated using experimental extraction curves (with simultaneous degradation) combined with experimental degradation curves, as schematically described in Fig. 2.5 below. The conclusion was that, theoretically, a 20 to 30% higher yield of anthocyanins could be obtained from red onions if there were no degradation problems. This would require an extraction time of around 30 min (plus 8 min pre-heating time). In another study (Co et al., 2009), an increasing extraction temperature resulted in higher antioxidant activity in both water and ethanol extracts of birch bark. Furthermore, the antioxidant activities did not level out at high extraction temperature; hence, it appeared that some types of reactions occurred, and that these resulted in higher antioxidant activity. For instance, hydrolysis of polyphenolic glycosides leads to aglycons, which results in compounds of higher antioxidant activity because of the higher number of ‘reactive’ hydroxyl groups. Hence, an important conclusion is that a higher temperature may result in higher extraction yield of the target molecules, but also degradative reactions as well as extraction of more unwanted compounds, i.e. a less selective extraction.
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52 Separation, extraction and concentration processes
25
Peak area (a.u.)
20
15
10
5
0
0
20
40 60 Time (min)
80
100
Fig. 2.5 Calculation of an ideal extraction curve (continuous lines on top) from experimentally obtained data on degradation (dashed line) and extraction with degradation (continuous line with circles), respectively, for cyanidin-3-(6≤ malonoylglucoside) in red onion (Petersson et al., 2010).
Another type of degradative reaction that may occur in food and agricultural samples is caramelization of sugars. When water is used as a solvent at temperatures of around 160 °C and above, sugars such as glucose caramelize (Montilla et al., 2006). Depending on the length of the extraction time, this obstacle may cause erratic results as well as plugging of tubings and filters. If such high temperatures are to be used, it is advisable to carry out the extraction as fast as possible, to minimize the occurance of caramelization. Maillard reactions may also occur during PFE, involving reducing sugars and amino acids or proteins, but also flavonoids, ascorbic acid and other carbonyl compounds (Manzocco et al., 2000). Maillard polymers can be a major contributor to antioxidant capacity at higher temperature (>170 °C), when using for instance water or water/ethanol mixtures as extraction solvents in food applications (Howard and Pandjaitan, 2008). Hence, in food applications, to be on the safe side, it is advisable to use an extraction temperature of less than 160 °C. 2.2.5 Extraction time The third most important parameter after extraction solvent and temperature, and maybe the absolutely most interesting one, is extraction time. Several studies have shown that PFE offers faster extraction methods relative to those based on Soxhlet, sonication and SFE, and similar to MAE (Dean, 1996;
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Principles of pressurized fluid extraction and applications 53 Focant et al., 2004; Vandenburg et al., 1999). For instance, a comparison study was made for the extraction of PAHs from contaminated soil using Soxhlet, two different MAE methods, SFE and PFE (Saim et al., 1997). The Soxhlet method needed 24 h, the SFE method 35 min, the two MAE methods 20 min and the PFE method only 10 min, and Soxhlet extraction gave the highest recovery and the worst precision. In another study, Heemken et al. (1997) made a comparison of PFE and SFE with Soxhlet extraction of PAHs, aliphatic hydrocarbons and chlorinated hydrocarbons from marine samples. Statistical evaluations of accuracy and precision showed that equivalent results were achieved. The extraction time needed for the PFE method was 15 min, whereas SFE needed 90 min and Soxhlet 24 h. In general, during method development, extraction time and temperature are the most important parameters that need to be optimized once the solvent or solvent mixture has been chosen, as well as the mode of sample pre-treatment. Typically, extraction times are in the range of 10–15 min, plus a preheating step of 5–10 min depending on the temperature used and instrumental capabilities. However, if a dynamic extraction system is used (Section 2.5.2), the solvent can be preheated and continuously pumped through the sample without needing to preheat the cell. In commercially available PFE systems, dynamic extraction is mimicked using several short cycles of static extraction, replacing a certain volume of extraction solvent in between each cycle. The most recent Dionex PFE system, ASE-350®, can accomplish semi-dynamic extractions by step-wise replacement of small volumes of solvent continuously during the extraction. 2.2.6 Extraction pressure An increased pressure in the PFE technique is mainly applied to keep the solvents as liquids, at temperatures above their atmospheric boiling point. Most PFE studies observe no difference in extraction yield when the pressure is varied in the range of 34–204 bar (Lundstedt et al., 2000), although it has been reported (Richter et al., 1996) that a higher pressure rendered a higher extraction yield, when a standard mixture of PAH and polychlorinated hydrocarbons from a reference material was spiked onto wet silica of different pore sizes, using methylene chloride/acetone (1:1 v/v) at 100 °C as a solvent. Higher pressure gave higher extraction yield, when the pore size of the wet silica was 300 Å, but no difference in recovery was observed with the dry silica. It was suggested that higher pressure could probably facilitate the extraction of analytes trapped inside matrix pores because the pressure would force the solvent into the pores of the matrix that normally would not be contacted by solvents at atmospheric pressure. Pressurized flow would also aid in the solubilization of air bubbles so that the solvent could more easily penetrate the sample matrix. Kremer et al. (2004) also obtained higher yield of acidic herbicides in soil when pressure was increased from 69 to 138 bar, at 100 °C with dichloromethane as solvent.
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54 Separation, extraction and concentration processes 2.2.7 Selectivity during the extraction Selective extractions are in general obtained by changing the solvent or the solvent power. Under ambient conditions (room temperature and atmospheric pressure) the solvent power of a liquid is essentially constant regardless of extraction conditions. In PFE, selectivity can mainly be obtained by varying the nature of the solvent or the solvent mixture (Majors, 1999; Ramos et al., 2002). It has been pointed out that a potential disadvantage with PFE is that the extraction tends to be exhaustive, therefore leading to nonselective extractions (Reighard and Olesik, 1996). Hawthorne et al. (2000) compared the difference in color of the extracts, when soils contaminated with PAH were extracted by Soxhlet, SFE, PFE and pressurized hot water extraction (PHWE). The Soxhlet and the PFE extracts were much darker mainly owing to the greater amount of extracted compounds, i.e. the extractions were less selective. Hence, to avoid coextracting compounds with maintained high yield of the target analytes, it is essential to choose the most selective solvent for the analyte combined with an appropriate temperature, i.e. high enough for efficient extraction of the target compounds, but not too high to avoid coextraction of unwanted molecules. Extraction time is also an important factor: the minimum time necessary should be used to achieve targeted extraction rather than exhaustive extraction of other matrix solutes. However, by using adsorbents in the extraction cell as an in-line clean-up step, more selective PFE extractions can be obtained. For instance, activated acidic alumina has been employed to adsorb fat when PCBs were extracted from spiked freeze-dried fish tissue, and in this way better chromatogram separations were achieved (Ezzell et al., 1996). Sulfuric acid impregnated silica has similarly been used to adsorb the lipids from fat-containing food and feed when extracting PCBs to obtain lipid-free extracts (Sporring and Björklund, 2004). In our own study (Nilsson et al., 2001), when extracting chlorinated paraffins from household waste, clean extracts and chromatograms were obtained when the drying agents Hydromatrix and sodium sulfate were used, a result which is in agreement with other studies (Hawthorne et al., 2000; Ramos et al., 2002). On the other hand, the use of adsorbents such as XAD-2, XAD-4, and XAD-16 did not provide the removal of unidentified interfering compounds. In another work of ours, squalene and a-tocopherol have been selectively extracted from olive oil without coextraction of triglycerides when utilizing Amberlite XAD-16 as adsorbent (data not published). A liquid chromatogram of the olive oil before and after PFE is illustrated in Fig. 2.6. 2.2.8 Accuracy and precision Accuracy and precision obtained with the various PFE methods are similar to those accomplished with other extraction techniques such as Soxhlet, sonication, SFE and MAE (Helaleh et al., 2005; Saim et al., 1997). As with any extraction technique, accuracy and precision of the chemical analysis
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Principles of pressurized fluid extraction and applications 55
2.5¥105
Squalene
2.0¥105 a-Tocopherol
UV detection (abs. units)
3.0¥105
1.5¥105 1.0¥105 5.0¥104 0
10.00
20.00
30.00 40.00 Time (min) (a)
50.00
60.00
30.00 40.00 Time (min) (b)
50.00
60.00
2.5¥105
Squalene
2.0¥105
a-Tocopherol
UV detection (abs. units)
3.0¥105
1.5¥105 1.0¥105 5.0¥104 0
10.00
20.00
Fig. 2.6 Liquid chromatograms of (a) olive oil diluted with ethanol/acetone and (b) a PFE extract of olive oil, obtained using XAD-16 as adsorbent and methanol as extraction solvent.
method heavily depend on how well the sample is homogenized, and if the collected sub-sample is an accurate representation of the entire sample under study. These problems with sample preparation can of course be minimized by taking out a larger number of sub-samples, i.e. replicates, for analysis. Another important variable is the particle size, i.e. too large a particle size leads to inefficient extraction and inaccurate results. This was demonstrated by Björklund et al. (1999) for sediment samples of different particle size distributions. Similarly, Isaac et al. (2005) in a study of the extraction of lipids from homogenized cod tissue obtained a higher yield compared with extracting intact wet cod muscle. The conclusion was that the proteins in the sample denaturated to a hard pellet during the extraction making the diffusion of the solvent into the matrix not efficient enough to extract all the lipids. On the other hand, too small a particle size led to a relatively compact sample, which in turn may result in channeling inside the sample and again inefficient extraction.
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56 Separation, extraction and concentration processes
2.3 Applications of pressurized fluid extraction Most of the studies on PFE deal with either environmental analysis or food and agricultural samples, targeting either pollutants or health-beneficial compounds such as antioxidants. Figure 2.7 shows the publication record from 1995 to 2008, taken from Web of Science. There is not enough space in this chapter to discuss in detail the wide range of different applications for PFE, hence, the reader is referred to a number of reviews of the field: on applications for environmental analysis (Giergielewicz-Mozajska et al., 2001; Nieto et al., 2008); food and drugs (Beyer and Biziuk, 2008; Carabias-Martinez et al., 2005; Herrero et al., 2006; Mendiola et al., 2007); and medicinal plants, herbs and agricultural (Huie, 2002; Wang and Weller, 2006). 2.3.1 Environmental applications Pesticides were extracted from fruits and vegetables (Adou et al., 2001; Barriada-Pereira et al., 2007; Blasco et al., 2005; Herrera et al., 2002; Pihlstrom et al., 2002; Tanaka et al., 2007; Wennrich et al., 2001) and from agricultural crops and soils (Hildebrandt et al., 2007; Otake et al., 2008; Popp et al., 1997; Schreck et al., 2008) using solvents and solvent mixtures such as acetone, acetone/dichloromethane, acetone/hexane, methanol, hexane saturated with acetonitrile, and acetonitrile and water, at temperatures between 60 and 130 °C. For polyhalogenated persistant organic pollutants such as polychlorinated biphenyls (PCBs), polybrominated diphenyl ethers (PBDEs), polychlorinated dibenzodioxins (PCDDs) and polychlorinated dibenzofurans (PCDFs), it is more common to use hexane, toluene/acetone and hexane/ acetone, at temperatures between 100 and 150 °C (Antunes et al., 2008; Bjorklund et al., 1999; Brandl et al., 2006), but water has also been used at temperatures between 250 and 300 °C (Yang et al., 1995). Number of publications
250 200 150 100 50
08 20
06
05
07 20
20
03
04
20
20
02
20
00
99
01
20
20
20
98
19
96
97
19
19
19
19
95
0 Year
Fig. 2.7 Publication record (Web of Science 2009-08-02, search term: ‘pressurized liquid extraction’ OR ‘accelerated solvent extraction’ OR ‘pressurized solvent extraction’ OR ‘subcritical water extraction’ OR ‘pressurized fluid extraction’). In total 1469 publications between 1995 and 2008.
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Principles of pressurized fluid extraction and applications 57 Organometallic compounds containing tin, arsenic and mercury have been extracted from sediments and soils (Beichert et al., 2000; Chiron et al., 2000), vegetables (Marcic et al., 2005) and seafood (Mato-Fernandez et al., 2007; Wahlen and Catterick, 2004) using methanol, water, methanol/water and methanol/ethyl acetate mixtures, at temperatures between ambient and 160 °C. There are also a few studies on the extraction of explosives from soil samples (Ragnvaldsson et al., 2007) and biological samples (Pan et al., 2005); and antibiotics from soil (Jacobsen et al., 2004; Schlusener et al., 2003) and meat samples (Berrada et al., 2008). Pharmaceutical applications are less common than environmental ones, and most of them consider the extraction of pharmaceutical drugs and their metabolites in environmental samples such as sediments, sludge and fish/ shell fish. For instance, pharmaceutical residues have been extracted from sludge using methanol/water (1:1) at 60 °C as solvent (Barron et al., 2008). There is also an increasing number of studies on endocrine disruptors in the environment, for instance, 17-b-estradiol has been extracted from soil using acetone/hexane (1:1) at 100 °C (Chun et al., 2005); and alkylphenolic compounds have been extracted from river sediment using methanol/acetone (1:1) at 50 °C as a solvent (Petrovic et al., 2002). 2.3.2 Food and agricultural applications Food and agricultural applications considering endogeneous compounds commonly involve the use of more environmentally friendly solvents such as water and ethanol, because these applications often aim at isolating healthbeneficial compounds for later use as food additives, or alternatively, the compounds that are extracted are fairly polar, hence stronger organic solvents are not needed. Exceptions are in applications concerning the extraction of lipids such as acylglycerols, sterols, terpenoids and essential oils, for which stronger organic solvents are generally employed, e.g. isopropanol, ethyl acetate and hexane. Flavonoids such as quercetin, kaempferol, catechin and anthocyanidins have been extracted from vegetables and herbal plants using pressurized hot water, ethanol or water/ethanol mixtures as solvents at temperatures between 50 and 160 °C (Arapitsas and Turner, 2008; Howard and Pandjaitan, 2008; Ibanez et al., 2003; Ollanketo et al., 2002; Turner et al., 2006). For instance, water and a 70:30 mixture of ethanol and water at temperatures of 50 to 190 °C were used to extract flavonoids from dried spinach (Howard and Pandjaitan, 2008). The results showed that the total phenolic content as well as the antioxidant capacity were the highest in extracts obtained with the highest extraction temperatures (170–190 °C) for both water and the water/ethanol mixture, although the solvent mixture was somewhat more efficient than pure water in extracting flavonoids. However, when separating the extracts into low (<1000 Da) and high (>1000 Da) molecular weight fractions, it turned out that at temperatures above 130 °C for water and
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58 Separation, extraction and concentration processes above 150 °C for ethanolic solvents, polymeric Maillard reaction products were responsible for the increase in antioxidant capacity. Howard and Pandjaitan (2008) concluded that, in order to extract flavonoids effectively, water should be used at temperatures between 50–130 °C and for ethanolic solvents, 50–150 °C. In one of our own studies, quercetin and its glycosides as well as isorhamnetin and kaempferol were extracted from yellow and red onion using water at 120 °C and 50 bar, employing three consecutive 5-min extraction cycles (Turner et al., 2006). Thermostable b-glucosidase was used to convert the polyphenolic glucosides into their respective aglycons within only 10 min of reaction, using water at 90 °C and pH 5 as a reaction media. Results obtained compared well with a conventional extraction method employing a 1:1 mixture of methanol and 2.4 M HCl at 80 °C and 2 h combined extraction/ reaction, followed by filtration. Phenolic compounds such as rosmarinic and carnosic acids, carnosol and methyl carnosate have been extracted by PHWE, ultrasonication-assisted methanol extraction, hydrodistillation, and maceration with 70% ethanol from sage (Ollanketo et al., 2002). It turned out that PHWE, performed at 100 °C, gave the highest antioxidant activity of the extracts. Furthermore, PHWE was faster than the conventional techniques used in this work. In another study (Ibanez et al., 2003), antioxidative compounds such as carnosol, rosmanol, carnosic acid, methyl carnosate, cirsimaritin and genkwanin have been extracted from rosemary leaves using subcritical water at temperatures between 25 and 200 °C, showing that the selectivity of the extraction can be tuned by varying the temperature. At 25 °C a high concentration of rosmanol was obtained in the extract, while at 200 °C, the dominant compound was carnosic acid. The antioxidant activity was high in all of the extracts obtained at 100, 150 and 200 °C, respectively (Ibanez et al., 2003). Phenolic acids have been extracted from fruits and vegetables using water, methanol and water/methanol mixtures, at temperatures ranging between 20 and 100 °C (Alonso-Salces et al., 2001; Chen et al., 2007a; Mukhopadhyay et al., 2006; Waksmundzka-Hainos et al., 2007). In general, the main concern for the extraction of phenolic acids as well as many of the flavanoids is the thermal stability, or rather the tendency to degrade during the extraction, as discussed in Section 2.2.4. Terpenoids including limonene, pinene, artemisinin, retinol, taxol, squalene, lycopene and carotene, have only limited water solubility even at elevated temperatures, hence these are extracted with, for instance, 2-propanol or methanol/ethyl acetate/light petroleum (1:1:1) at temperatures ranging from 40 to 190 °C (Breithaupt, 2004; Fojtova et al., 2008; Schaneberg and Khan, 2002; Waldeback et al., 2004). There are, however, also studies in which pressurized hot water is used as a solvent; Yang et al. (2007) used water at temperatures between 100 and 250 °C to extract a-pinene, limonene, camphor, citronellol, and carvacrol terpenoids from oregano and basil leaves. However, it was shown that thermal degradation occurred at all temperatures tested,
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Principles of pressurized fluid extraction and applications 59 although least severe at 100 °C – between 10 and 25% for a 30-min treatment in static mode. In a work by Breithaupt (2004), carotenoids were extracted by PFE and compared with a conventional solvent extraction, in both cases using a solvent mixture of methanol/ethyl acetate/light petroleum (1:1:1). The PFE extraction conditions were optimized by varying the temperature (25–80 °C) and the pressure (70–140 bar), giving the highest yield at 40 °C and 70 bar. Both methods were used for the extraction of a large number of different carotenoids from beverages, pudding mixes, cereals, cookies and sausages (Breithaupt, 2004). Acylglycerols and sterols have been extracted from eggs and egg-containing food products, using chloroform/methanol 2:1 (v/v) and hexane/isopropanol 3:2 (v/v), at various extraction temperatures and pressures (60 °C and 150 bar, 100 °C at 150 bar, and 120 °C and 200 bar (Boselli et al., 2001). Results showed that the hexane/isopropanol mixture employed at 60 °C and 150 bar could successfully be used to extract lipids from egg products. In another study, lipids were extracted from freshly ground corn kernels and ground rolled oats, using four different organic solvents: hexane, dichloromethane, isopropanol, and ethanol, at two temperatures, 40 and 100 °C (Moreau et al., 2003). The results showed that oat oil with the highest levels of digalactosyldiacylglycerol (DGDG) was obtained using ethanol at 100 °C. Comparison of the extraction effects with seeds of two different species of grains indicated that the extractability is greatly dependent on the polarity and temperature of the solvent. Tocopherols and tocotrienols are fat-soluble compounds extractable using solvents such as methanol and acetone. For instance, tocopherols and tocotrienols have been extracted from cereals using methanol as extraction solvent at 50 °C and 110 bar, using only one 5-min extraction cycle (Bustamante-Rangel et al., 2007). Other food and agricultural applications include for instance alkaloids (Mroczek and Mazurek, 2009); saponins (Chen et al., 2007b); coumarins (Waksmundzka-Hajnos et al., 2004); oligosaccharides (Bansleben et al., 2008); lignans (Smeds et al., 2007); and various essential oils (JimenezCarmona et al., 1999; Ozel et al., 2003; Schaneberg and Khan, 2002).
2.4 Future trends The authors believe that some of the most important future trends for PFE are: (i) switching to more environmentally friendly solvents; (ii) hyphenation and automation; (iii) miniaturization; (iv) mimic nature; and (v) transfer to industry.
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60 Separation, extraction and concentration processes All these trends are somewhat connected with the growing push for a more sustainable development. In order to minimize the use of organic solvents that need to be produced and then disposed of, both switching to more environmentally friendly solvents and miniaturizing the extraction equipment is vital. The use of solvents at elevated temperatures implies that more efficient extractions can be obtained, even when using more environmentally friendly solvents such as water and ethanol rather than the classical ones such as hexane and ethyl acetate. Hence, future trends will most probably involve switching from conventional organic solvent based extraction methods to PFE using water, ethanol and other ‘green’ alternatives. This has already been demonstrated (Burkhardt et al., 2005; Co et al., 2009; Crescenzi et al., 1999; Curren and King, 2001; Fernandez-Perez and de Castro, 2000; Hawthorne et al., 1998; Hyotylainen et al., 2000; Morales-Munoz et al., 2002; Smith, 2002; Turner et al., 2006; Yang and Li, 1999), but more applications are expected in the near future. Regarding miniaturization of PFE, we found only one example (Ramos et al., 2000), in which PAHs were extracted from soil and sediment using an extraction vessel constructed by a 10 mm ¥ 3 mm inner diameter stainlesssteel tubing. The only extraction solvent used was 100 mL of toluene at 200 °C and 150 bar, in an extraction process taking 10 min, with subsequent direct injection into large-volume injection gas chromatography–mass spectrometry (GC–MS). Future trends will probably include more examples of miniaturized PFE systems, probably also combined with analytical separation and detection systems. It has been recognized that in order to minimize the production of waste and emissions of carbon dioxide, it is advantageous to use as few steps in a process as possible, i.e. combined (hyphenated) methods and automation are relevant. There are several good examples showing that PFE can be coupled on-line with separation by HPLC (Hyotylainen et al., 2000; Li et al., 2000), but also with GC using a miniaturized PFE system (Ramos et al., 2000). Smith showed that subcritical water extraction can be coupled on-line with chromatography using subcritical water as the only mobile phase (Smith, 2002; Tajuddin and Smith, 2002). In the future, technology development enabling combination with several other separation and detection techniques is expected, as well as in situ analysis instrumentation. In order to better understand processes occurring in real life or in the environment, laboratory-scale methods that mimic nature in an efficient way are important. Thus, by using PFE, an interesting possibility is to simulate what could happen in a real life situation (in the human body, in the environment or within an industrial process). This can be accomplished using a suitable solvent, similar to the surroundings in question, for example water with appropriate pH and ion concentrations (such as seawater). Anaerobic conditions can be simulated using a nitrogen purge between sequential extractions. Hence, it is possible to predict what would happen if a specific chemical is used by a specific company in a specific environment
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Principles of pressurized fluid extraction and applications 61 and, in this way, obtain detailed information for an environmental regulatory impact analysis. Another challenge would be to follow the migration of any hazardous compound and their degradation products in soil, sediment, sludge or biota. Ongoing research in this area is expected to address important environmental issues. A final note on future trends of PFE is that several of the smaller-scale processes will be transferred to industry to replace conventional solvent extraction methods in chemical analysis applications, as well as scaled up for industrial processes in applications regarding isolation of interesting (valuable) compounds from the plant kingdom. For instance, a standard method (US EPA SW-846 Method 3545) for the extraction of PCBs from sediments using hexane/acetone (1:1) as a solvent at 150 °C has replaced a traditional Soxhlet method. Both new applications and replacement of technology in old applications are expected to be developed in the near future.
2.5 Sources of further information and advice 2.5.1 Commercially available equipment There is equipment available on the market that performs static PFE, with full or semi-automation. Currently, the following vendors have been identified: Dionex (www.dionex.com) products include ASE® 150 (one extraction vessel at the time, of sizes 1, 5, 10, 22, 34, 66 or 100 mL) and ASE® 350 (carousel with up to 24 extraction vessels of the same sizes as for the 150-system). A solvent controller allows mixing and delivery of up to three solvents. Temperature range applicable is 40–200 °C and the pressure is fixed to 1500 psi (103 bar). pH hardened pathway with Dionium™ components and extraction vessels make the instrument compatible with acid or alkaline pretreated sample matrices. ∑ Applied Separations (www.appliedseparations.com) has two extraction systems, one PSE™ and fast-PSE™, which operate with one and six in-parallel extraction vessels, respectively. Available vessel sizes are 11, 22 and 33 mL. An automated solvent dispenser is optional, and up to four different solvents can be handled. Temperature range is 50– 200 °C and maximum pressure is 150 bar. ∑ Fluid Management Systems (www.fmsenvironmental.com) market PFE™ systems that handle one to six samples (six modules) at the time in parallel, with the option of in-line clean up by column chromatography. Temperature and pressure ranges obtainable are 70–200 °C and 1500– 3000 psi (103–207 bars), respectively. Extraction vessels come in sizes ranging between 5–250 mL. ∑
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62 Separation, extraction and concentration processes 2.5.2 Designing your own equipment In our laboratory, we usually build our own equipment, since the ones commercially available are all based on static PFE and, in addition, operate within a strictly limited temperature range. Hence, in order to run dynamic PFE, even at temperatures above 200 °C, it is necessary to custom build PFE instrumentation. What we recommend is to use the following parts: (a) an optional heater plate set to a temperature of around 20–30 °C below the boiling point of the solvent; (b) a bottle containing the extraction solvent or solvent mixture; (c) an ordinary HPLC pump; (d) a long stainless-steel tubing that enables the solvent to reach desired temperature before passing through the extraction vessel; (e) a type K thermocouple temperature probe for measuring the temperature of the solvent just before entering the extraction vessel; (f) a stainless-steel extraction vessel (for instance a preparative HPLC column) in which the sample is placed before extraction; (g) a simple GC oven (only temperature control is needed, hence no need for injector or detector); (h) a cooling bath to bring down the temperature of the extractant to below the boiling point of the solvent; (i) a needle valve or a back pressure regulator; (j) collection vial(s); and (k) optional nitrogen gas to flush out the entire system after extraction. It is advisable to install an in-line filter to protect the needle valve/back pressure regulator. It is also recommended to install a pressure gauge for control of the pressure as well as a burst disc in case the pressure rises too high in the system. There are burst discs available for different pressures and solvents. Figure 2.8 shows a schematic of a home built dynamic PFE system. (c)
(b)
VALCO (a)
(k)
(e)
(g) (d)
(i)
(f) (h) (j)
Fig. 2.8 Schematic of a home-built dynamic PFE system. (a) Heater plate; (b) extraction solvent; (c) HPLC pump; (d) stainless steel tubing; (e) temperature probe; (f) extraction vessel; (g) oven; (h) cooling bath; (i) needle valve; (j) collection vial; (k) nitrogen gas.
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Principles of pressurized fluid extraction and applications 63 A factor to consider is which solvents and temperatures are allowable by the different parts of the extraction system. For instance, an HPLC pump cannot take too high a temperature, and especially not with some of the organic solvents. If water is used as a solvent, caution must be taken if temperatures above ca. 200 °C are used, since water becomes more and more corrosive at higher temperature. Hence, higher-quality nickel steel alloy, such as Hastelloy® that exhibits high resistance to corrosion, should be used. However, owing to obstacles such as caramelization and Maillard reactions occurring in food and agricultural samples, temperatures applied are rarely above 160 °C. Finally, if an organic solvent is used, safety arrangements around your homebuilt device, such as ventilation and burst discs connected to stainless-steel tubings leading to a secure waste container, should be provided.
2.6 Conclusions In industrial and governmental analytical and research laboratories, PFE is a promising extraction technique that easily replaces conventional techniques such as Soxhlet extraction – especially since the same solvents can be used. Compared with conventional extraction techniques, PFE is faster; more automated; and enables a ‘friendlier’ work environment. Furthermore, new PFE methods are fairly easy to develop by optimizing temperature and extraction time. Since PFE operates at elevated temperatures, it is possible to use less harmful solvents than in Soxhlet with maintained extraction efficiency. However, caution has to be taken regarding degradation of thermolabile target compounds during the extraction, even though the problems are usually minor if the extraction time and temperature are carefully optimized. Finally, selectivity can be obtained in PFE by firstly selecting an appropriate extraction solvent, secondly by varying the temperature of the solvent, which is especially effective for water, or thirdly by using adsorbents in the extraction cell. How does PFE fit into a sustainable development? PFE uses smaller amounts of organic solvent, thereby preventing unnecessary waste, because less organic solvent needs to be manufactured and less waste is thereby produced by the manufacturer. In addition, less organic solvent waste is produced at the laboratory and, in total, there will be a decrease in volatile organic carbon emissions and a reduced risk for photochemical smog formation. Accidental emissions to municipal wastewater plants will do less harm. By being able to substitute hazardous solvents with more environmentally friendly solvents such as ethanol and water, the toxicity is reduced with positive effects on both the environment and health in general. Creating high diffusion liquids requires a significant amount of energy, which is a disadvantage. However, because the extraction times are shorter, minutes compared with hours in Soxhlet and modified Soxhlet techniques, and, as these traditional solvent extraction techniques are usually performed © Woodhead Publishing Limited, 2010
64 Separation, extraction and concentration processes just below the boiling point, a significant amount of energy is saved. Lifecycle assessment should be conducted to calculate the overall energy usage in different methodologies.
2.7 References Adou, K., Bontoyan, W. R. and Sweeney, P. J. (2001), ‘Multiresidue method for the analysis of pesticide residues in fruits and vegetables by accelerated solvent extraction and capillary gas chromatography’, J. Agric. Food Chem., 49, 4153–60. Alonso-Salces, R. M., Korta, E., Barranco, A., Berrueta, L. A., Gallo, B. and Vicente, F. (2001), ‘Determination of polyphenolic profiles of Basque cider apple varieties using accelerated solvent extraction’, J. Agric. Food Chem., 49, 3761–67. Antunes, P., Viana, P., Vinhas, T., Capelo, J. L., Rivera, J. and Gaspar, E. (2008), ‘Optimization of pressurized liquid extraction (PFE) of dioxin–furans and dioxin-like PCBs from environmental samples’, Talanta, 75, 916–25. Arapitsas, P. and Turner, C. (2008), ‘Pressurized solvent extraction and monolithic columnHPLC/DAD analysis of anthocyanins in red cabbage’, Talanta, 74, 1218–23. Bansleben, D., Schellenberg, I. and Wolff, A. C. (2008), ‘Highly automated and fast determination of raffinose family oligosaccharides in Lupinus seeds using pressurized liquid extraction and high-performance anion-exchange chromatography with pulsed amperometric detection’, J. Sci. Food Agric., 88, 1949–53. Barriada-Pereira, M., Gonzalez-Castro, M. J., Muniategui-Lorenzo, S., Lopez-Mahia, P., Prada-Rodriguez, D. and Fernandez-Fernandez, E. (2007), ‘Comparison of pressurized liquid extraction and microwave assisted extraction for the determination of organochlorine pesticides in vegetables’, Talanta, 71, 1345–51. Barron, L., Tobin, J. and Paull, B. (2008), ‘Multi-residue determination of pharmaceuticals in sludge and sludge enriched soils using pressurized liquid extraction, solid phase extraction and liquid chromatography with tandem mass spectrometry’, J. Environ. Monit., 10, 353–61. Bautz, H., Polzer, J. and Stieglitz, L. (1998), ‘Comparison of pressurised liquid extraction with Soxhlet extraction for the analysis of polychlorinated dibenzo-p-dioxins and dibenzofurans from fly ash and environmental matrices’, J. Chromatogr A, 815, 231–41. Beichert, A., Padberg, S. and Wenclawiak, B. W. (2000), ‘Selective determination of alkylmercury compounds in solid matrices after subcritical water extraction, followed by solid-phase microextraction and GC–MS’, Appl. Organometal. Chem., 14, 493–98. Berrada, H., Borrull, F., Font, G. and Marce, R. M. (2008), ‘Determination of macrolide antibiotics in meat and fish using pressurized liquid extraction and liquid chromatography–mass spectrometry’, J. Chromatogr. A, 1208, 83–89. Beyer, A. and Biziuk, M. (2008), ‘Applications of sample preparation techniques in the analysis of pesticides and PCBs in food’, Food Chem., 108, 669–80. Bjorklund, E., Bowadt, S., Nilsson, T. and Mathiasson, L. (1999), ‘Pressurized fluid extraction of polychlorinated biphenyls in solid environmental samples’, J. Chromatogr. A, 836, 285–93. Bjorklund, E., Turner, C., Karlsson, L., Mathiasson, L., Sivik, B. and Skogsmo, J. (1996), ‘The influence of oil extractability and metal part geometry in degreasing processes using supercritical carbon dioxide’, J. Supercrit Fluids, 9, 56–60. Blasco, C., Font, G. and Pico, Y. (2005), ‘Analysis of pesticides in fruits by pressurized liquid extraction and liquid chromatography-ion trap-triple stage mass spectrometry’, J. Chromatogr. A, 1098, 37–43.
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Principles of pressurized fluid extraction and applications 65 Boselli, E., Velazco, V., Caboni, M. F. and Lercker, G. (2001), ‘Pressurized liquid extraction of lipids for the determination of oxysterols in egg-containing food’, J. Chromatogr. A, 917, 239–44. Brandl, R. C., Bucheli, T. D., Kupper, T., Stadelmann, F. X. and Tarradellas, J. (2006), ‘Optimised accelerated solvent extraction of PCBs and PAHs from compost’, Int. J. Environ. Anal. Chem., 86, 505–25. Breithaupt, D. E. (2004), ‘Simultaneous HPLC determination of carotenoids used as food coloring additives: applicability of accelerated solvent extraction’, Food Chem., 86, 449–56. Burkhardt, M. R., Zaugg, S. D., Burbank, T. L., Olson, M. C. and Iverson, J. L. (2005), ‘Pressurized liquid extraction using water/isopropanol coupled with solid-phase extraction cleanup for semivolatile organic compounds, polycyclic aromatic hydrocarbons (PAH), and alkylated PAH homolog groups in sediment’, Anal. Chim. Acta, 549, 104–16. Bustamante-Rangel, M., Delgado-Zamarreno, M. M., Sanchez-Perez, A. and CarabiasMartinez, R. (2007), ‘Determination of tocopherols and tocotrienols in cereals by pressurized liquid extraction-liquid chromatography-mass spectrometry’, Anal. Chim. Acta, 587, 216–21. Carabias-Martinez, R., Rodriguez-Gonzalo, E., Revilla-Ruiz, P. and Hernandez-Mendez, J. (2005), ‘Pressurized liquid extraction in the analysis of food and biological samples’, J. Chromatogr. A, 1089, 1–17. Chen, C. R., Lee, Y. N., Chang, C. M. J., Lee, M. R. and Wei, I. C. (2007a), ‘Hotpressurized fluid extraction of flavonoids and phenolic acids from Brazilian propolis and their cytotoxic assay in vitro’, J. Chin. Inst. Chem. Eng., 38, 191–96. Chen, J. H., Li, W. L., Yang, B. J., Guo, X. C., Lee, F. S. C. and Wang, X. O. (2007b), ‘Determination of four major saponins in the seeds of Aesculus chinensis Bunge using accelerated solvent extraction followed by high-performance liquid chromatography and electrospray-time of flight mass spectrometry’, Anal. Chim. Acta, 596, 273–80. Chiron, S., Roy, S., Cottier, R. and Jeannot, R. (2000), ‘Speciation of butyl- and phenyltin compounds in sediments using pressurized liquid extraction and liquid chromatographyinductively coupled plasma mass spectrometry’, J. Chromatogr. A, 879, 137–45. Chun, S., Lee, J., Geyer, R. and White, D. C. (2005), ‘Comparison of three extraction methods for 17 b-estradiol in sand, bentonite, and organic-rich silt loam’, J. Environ. Sci. Health Part B Pestic. Food Contam. Agric. Wastes, 40, 731–40. Co, M., Koskela, P., Eklund-Åkergren, P., Srinivas, K., King, J. W., Sjöberg, P. J. R. and Turner, C. (2009), ‘Pressurized liquid extraction of betulin and antioxidants from birch bark’, Green Chem., 11, 668–74. Cotton, N. J., Bartle, K. D., Clifford, A. A. and Dowle, C. J. (1993), ‘Rate and extent of supercritical fluid extraction of additives from polypropylene: diffusion, solubility, and matrix effects’, J. Appl. Polymer Sci., 48, 1607–19. Crescenzi, C., D’Ascenzo, G., Di Corcia, A., Nazzari, M., Marchese, S. and Samperi, R. (1999), ‘Multiresidue herbicide analysis in soil: Subcritical water extraction with an on-line sorbent trap’, Anal. Chem., 71, 2157–63. Curren, M. S. S. and King, J. W. (2001), ‘Ethanol-modified subcritical water extraction combined with solid-phase microextraction for determining atrazine in beef kidney’, J. Agric. Food Chem., 49, 2175–80. Dean, J. R. (1996), ‘Accelerated solvent extraction of polycyclic aromatic hydrocarbons from contaminated soil’, Anal. Commun., 33, 191–92. Ezzell, J., Richter, B. and Francis, E. (1996), ‘Selective extraction of polychlorinated biphenyls from fish tissue using accelerated solvent extraction’, Am. Environ. Lab, 12, 12–13. Fernandez-Perez, V. and de Castro, M. D. L. (2000), ‘Micelle formation for improvement of continuous subcritical water extraction of polycyclic aromatic hydrocarbons in soil prior to high-performance liquid chromatography-fluorescence detection’, J. Chromatogr. A, 902, 357–67.
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66 Separation, extraction and concentration processes Fitzpatrick, L. J. and Dean, J. R. (2002), ‘Extraction solvent selection in environmental analysis’, Anal. Chem., 74, 74–79. Focant, J. F., Pirard, C. and De Pauw, E. (2004), ‘Automated sample preparation –fractionation for the measurement of dioxins and related compounds in biological matrices: a review’, Talanta, 63, 1101–13. Fojtova, J., Lojkova, L. and Kuban, V. (2008), ‘GC/MS of terpenes in walnut-tree leaves after accelerated solvent extraction’, J. Sep. Sci., 31, 162–68. Giergielewicz-Mozajska, H., Dabrowski, L. and Namiesnik, J. (2001), ‘Accelerated Solvent Extraction (ASE) in the analysis of environmental solid samples – some aspects of theory and practice’, Crit. Rev. Anal. Chem., 31, 149–65. Hansen, C. M. (1967), ‘The three dimensional solubility parameter–key to paint component affinities: II’, J. Paint Technol., 511, 104–17. Hansen, C. M., 2007, Solubility parameters – user’s handbook, Boca Raton, Florida, CRC Press. Hawthorne, S. B., Yang, Y. and Miller, D. J. (1994), ‘Extraction of organic pollutants from environmental solids with subcritical and supercritical water’, Anal. Chem., 66, 2912–20. Hawthorne, S. B., Grabanski, C. B., Hageman, K. J. and Miller, D. J. (1998), ‘Simple method for estimating polychlorinated biphenyl concentrations on soils and sediments using subcritical water extraction coupled with solid-phase microextraction’, J. Chromatogr. A, 814, 151–60. Hawthorne, S. B., Grabanski, C. B., Martin, E. and Miller, D. J. (2000), ‘Comparisons of Soxhlet extraction, pressurized liquid extraction, supercritical fluid extraction and subcritical water extraction for environmental solids: recovery, selectivity and effects on sample matrix’, J. Chromatogr. A, 892, 421–33. Heemken, O. P., Theobald, N. and Wenclawiak, B. W. (1997), ‘Comparison of ASE and SFE with Soxhlet, sonication, and methanolic saponification extractions for the determination of organic micropollutants in marine particulate matter’, Anal. Chem., 69, 2171–80. Helaleh, M. I. H., Al-Omair, A., Ahmed, N. and Gevao, B. (2005), ‘Quantitative determination of organochlorine pesticides in sewage sludges using Soxtec, Soxhlet and pressurized liquid extractions and ion trap mass-mass spectrometric detection’, Anal. Bioanal. Chem., 382, 1127–34. Herrera, M. C., Prados-Rosales, R. C., Luque-Garcia, J. L. and de Castro, M. D. L. (2002), ‘Static–dynamic pressurized hot water extraction coupled to on-line filtration – solid-phase extraction – high-performance liquid chromatography – post-column derivatization – fluorescence detection for the analysis of N-methylcarbamates in foods’, Anal. Chim. Acta, 463, 189–97. Herrero, M., Cifuentes, A. and Ibanez, E. (2006), ‘Sub- and supercritical fluid extraction of functional ingredients from different natural sources: Plants, food-by-products, algae and microalgae – a review’, Food Chem., 98, 136–48. Hildebrandt, A., Lacorte, S. and Barcelo, D. (2007), ‘Assessment of priority pesticides, degradation products, and pesticide adjuvants in groundwaters and top soils from agricultural areas of the Ebro river basin’, Anal. Bioanal. Chem., 387, 1459–68. Hildebrand, J. H. and Scott, R. L., 1962, Regular solutions, Englewood Cliffs, NJ, Prentice-Hall Inc. Hildebrand, J. H. and Scott, R. L., 1964, The solubility of nonelectrolytes, 3rd Edition, New York, Dover Publication, Inc. Howard, L. and Pandjaitan, N. (2008), ‘Pressurized liquid extraction of flavonoids from spinach’, J. Food Sci., 73, C151–C57. Huie, C. W. (2002), ‘A review of modern sample-preparation techniques for the extraction and analysis of medicinal plants’, Anal. Bioanal. Chem., 373, 23–30. Hyotylainen, T., Andersson, T., Hartonen, K., Kuosmanen, K. and Riekkola, M. L. (2000), ‘Pressurized hot water extraction coupled on line with LC–GC: determination
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Principles of pressurized fluid extraction and applications 67 of polyaromatic hydrocarbons in sediment’, Anal. Chem., 72, 3070–76. Ibanez, E., Kubatova, A., Senorans, F. J., Cavero, S., Reglero, G. and Hawthorne, S. B. (2003), ‘Subcritical water extraction of antioxidant compounds from rosemary plants’, J. Agric. Food Chem., 51, 375–82. Isaac, G., Waldeback, M., Eriksson, U., Odham, G. and Markides, K. E. (2005), ‘Total lipid extraction of homogenized and intact lean fish muscles using pressurized fluid extraction and batch extraction techniques’, J. Agric. Food Chem., 53, 5506–12. Jacobsen, A. M., Halling-Sorensen, B., Ingerslev, F. and Hansen, S. H. (2004), ‘Simultaneous extraction of tetracycline, macrolide and sulfonamide antibiotics from agricultural soils using pressurised liquid extraction, followed by solid-phase extraction and liquid chromatography-tandem mass spectrometry’, J. Chromatogr. A, 1038, 157–70. Jimenez-Carmona, M. M., Ubera, J. L. and de Castro, M. D. L. (1999), ‘Comparison of continuous subcritical water extraction and hydrodistillation of marjoram essential oil’, J. Chromatogr. A, 855, 625–32. Kremer, E., Rompa, M. and Zygmunt, B. (2004), ‘Extraction of acidic herbicides from soil by means of accelerated solvent extraction’, Chromatographia, 60, S169–S74. Li, B., Yang, Y., Gan, Y. X., Eaton, C. D., He, P. and Jones, A. D. (2000), ‘On-line coupling of subcritical water extraction with high-performance liquid chromatography via solid-phase trapping’, J. Chromatogr. A, 873, 175–84. Lindahl, S., Ekman, A., Khan, S., Wennerberg, C., Börjesson, P., Sjöberg, P. J. R., Nordberg-Karlsson, E. and Turner, C. (2010), ‘Exploring the possibilities of using a thermostable mutant of b-glucosidase for rapid hydrolysis of quercetin glucosides in hot water’, Green Chem., 12, 159–168. Lou, X. W., Janssen, H. G. and Cramers, C. A. (1997), ‘Parameters affecting the accelerated solvent extraction of polymeric samples’, Anal. Chem., 69, 1598–603. Lundstedt, S., van Bavel, B., Haglund, P., Tysklind, M. and Oberg, L. (2000), ‘Pressurised liquid extraction of polycyclic aromatic hydrocarbons from contaminated soils’, J. Chromatogr. A, 883, 151–62. Majors, R. E. (1996), ‘The changing role of extraction in preparation of solid samples’, LC GC, 14, 88–96. Majors, R. E. (1999), ‘An overview of sample preparation methods for solids’, LC GC, 17, 8–13. Manzocco, L., Calligaris, S., Mastrocola, D., Nicoli, M. C. and Lerici, C. R. (2000), ‘Review of nonenzymatic browning and antioxidant capacity in processed foods’, Trends Food Sci. Technol., 11, 340–46. Marcic, C., Lespes, G. and Potin-Gautier, M. (2005), ‘Pressurised solvent extraction for organotin speciation in vegetable matrices’, Anal. Bioanal. Chem., 382, 1574–83. Mato-Fernandez, M. J., Otero-Rey, J. R., Moreda-Pineiro, J., Alonso-Rodriguez, E., Lopez-Mahia, P., Muniategui-Lorenzo, S. and Prada-Rodriguez, D. (2007), ‘Arsenic extraction in marine biological materials using pressurised liquid extraction’, Talanta, 71, 515–20. Mendiola, J. A., Herrero, M., Cifuentes, A. and Ibanez, E. (2007), ‘Use of compressed fluids for sample preparation: food applications’, J. Chromatogr. A, 1152, 234–46. Montilla, A., Ruiz-Matute, A. I., Sanz, M. L., Martìnez-Castro, I. and del Castillo, M. D. (2006), ‘Difructose anhydrides as quality markers of honey and coffee’, Food Res. Int., 39, 801–06. Morales-Munoz, S., Luque-Garcia, J. L. and de Castro, M. D. L. (2002), ‘Pressurized hot water extraction with on-line fluorescence monitoring: a comparison of the static, dynamic, and static-dynamic modes for the removal of polycyclic aromatic hydrocarbons from environmental solid samples’, Anal. Chem., 74, 4213–019. Moreau, R. A., Powell, M. J. and Singh, V. (2003), ‘Pressurized liquid extraction of polar and nonpolar lipids in corn and oats with hexane, methylene chloride, isopropanol, and ethanol’, J. Am. Oil Chem. Soc., 80, 1063–67. Mroczek, T. and Mazurek, J. (2009), ‘Pressurized liquid extraction and anticholinesterase
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68 Separation, extraction and concentration processes activity-based thin-layer chromatography with bioautography of Amaryllidaceae alkaloids’, Anal. Chim. Acta, 633, 188–96. Mukhopadhyay, S., Luthria, D. L. and Robbins, R. J. (2006), ‘Optimization of extraction process for phenolic acids from black cohosh (Cimicifuga racemosa) by pressurized liquid extraction’, J. Sci. Food Agric., 86, 156–62. Nemoto, S. and Lehotay, S. J. (1998), ‘Analysis of multiple herbicides in soybeans using pressurized liquid extraction and capillary electrophoresis’, J. Agric. Food. Chem., 46, 2190–99. Nieto, A., Borrull, F., Marce, R. M. and Pocurull, E. (2008), ‘Pressurized liquid extraction of contaminants from environmental samples’, Curr. Anal. Chem., 4, 157–67. Nilsson, M. L., Waldeback, M., Liljegren, G., Kylin, H. and Markides, K. E. (2001), ‘Pressurized-fluid extraction (PFE) of chlorinated paraffins from the biodegradable fraction of source-separated household waste’, Fresenius. J. Anal. Chem., 370, 913–18. Ollanketo, M., Peltoketo, A., Hartonen, K., Hiltunen, R. and Riekkola, M. L. (2002), ‘Extraction of sage (Salvia officinalis L.) by pressurized hot water and conventional methods: antioxidant activity of the extracts’, Eur. Food Res. Technol., 215, 158– 63. Otake, T., Aoyagi, Y., Matsuo, M., Itoh, N. and Yarita, T. (2008), ‘Evaluation of pressurized liquid extraction for the analysis of four pesticides in unpolished rice’, J. Environ. Sci. Health Part B Pestic. Food Contam. Agric. Wastes, 43, 390–94. Ozel, M. Z., Gogus, F. and Lewis, A. C. (2003), ‘Subcritical water extraction of essential oils from Thymbra spicata’, Food Chem., 82, 381–86. Pan, X. P., Zhang, B. H. and Cobb, G. P. (2005), ‘Extraction and analysis of trace amounts of cyclonite (RDX) and its nitroso-metabolites in animal liver tissue using gas chromatography with electron capture detection (GC–ECD)’, Talanta, 67, 816–23. Perry, R. H., Green, D. W. and Maloney, J. O., 1984, Perry’s chemical engineers handbook, New York, Mc Graw–Hill. Petersson, E. V., Puerta, A., Bergquist, J. and Turner, C. (2008), ‘Analysis of anthocyanins in red onion using capillary electrophoresis – time of flight-mass spectrometry’, Electrophoresis, 29, 2723–30. Petersson, E. V., Liu, J., Sjöberg, P. J., Danielsson, R. and C. Turner (2010), Pressurized hot water extraction of anthocyanins from red onion: a study on extraction and degradation rates, Anal. Chim Acta, 663, 27–32. Petrovic, M., Lacorte, S., Viana, P. and Barcelo, D. (2002), ‘Pressurized liquid extraction followed by liquid chromatography-mass spectrometry for the determination of alkylphenolic compounds in river sediment’, J. Chromatogr. A, 959, 15–23. Pihlstrom, T., Isaac, G., Waldeback, M., Osterdahl, B. G. and Markides, K. E. (2002), ‘Pressurised fluid extraction (PFE) as an alternative general method for the determination of pesticide residues in rape seed’, Analyst, 127, 554–59. Popp, P., Keil, P., Moder, M., Paschke, A. and Thuss, U. (1997), ‘Application of accelerated solvent extraction followed by gas chromatography, high-performance liquid chromatography and gas chromatography mass spectrometry for the determination of polycyclic aromatic hydrocarbons, chlorinated pesticides and polychlorinated dibenzop-dioxins and dibenzofurans in solid wastes’, J. Chromatogr. A, 774, 203–11. Ragnvaldsson, D., Brochu, S. and Wingfors, H. (2007), ‘Pressurized liquid extraction with water as a tool for chemical and toxicological screening of soil samples at army live-fire training ranges’, J. Hazard. Mater., 142, 418–24. Ramos, L., Vreuls, J. J. and Brinkman, U. A. T. (2000), ‘Miniaturised pressurised liquid extraction of polycyclic aromatic hydrocarbons from soil and sediment with subsequent large-volume injection-gas chromatography’, J. Chromatogr. A, 891, 275–86. Ramos, L., Kristenson, E. M. and Brinkman, U. A. T. (2002), ‘Current use of pressurised liquid extraction and subcritical water extraction in environmental analysis’, J. Chromatogr. A, 975, 3–29.
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Principles of pressurized fluid extraction and applications 69 Reighard, T. S. and Olesik, S. V. (1996), ‘Bridging the gap between supercritical fluid extraction and liquid extraction techniques: alternative approaches to the extraction of solid and liquid environmental matrices’, Crit. Rev. Anal. Chem., 26, 61–99. Richter, B. E., Jones, B. A., Ezzell, J. L., Porter, N. L., Avdalovic, N. and Pohl, C. (1996), ‘Accelerated solvent extraction: A technique for sample preparation’, Anal. Chem., 68, 1033–39. Saim, N., Dean, J. R., Abdullah, M. P. and Zakaria, Z. (1997), ‘Extraction of polycyclic aromatic hydrocarbons from contaminated soil using Soxhlet extraction, pressurised and atmospheric microwave-assisted extraction, supercritical fluid extraction and accelerated solvent extraction’, J. Chromatogr. A, 791, 361–66. Schaneberg, B. T. and Khan, I. A. (2002), ‘Comparison of extraction methods for marker compounds in the essential oil of lemon grass by GC’, J. Agric. Food Chem., 50, 1345–49. Schlusener, M. P., Spiteller, M. and Bester, K. (2003), ‘Determination of antibiotics from soil by pressurized liquid extraction and liquid chromatography-tandem mass spectrometry’, J. Chromatogr. A, 1003, 21–28. Schreck, E., Geret, F., Gontier, L. and Treilhou, M. (2008), ‘Development and validation of a rapid multiresidue method for pesticide determination using gas chromatographymass spectrometry: a realistic case in vineyard soils’, Talanta, 77, 298–303. Smeds, A. I., Eklund, P. C., Sjoholm, R. E., Willfor, S. M., Nishibe, S., Deyama, T. and Holmbom, B. R. (2007), ‘Quantification of a broad spectrum of lignans in cereals, oilseeds, and nuts’, J. Agric. Food Chem., 55, 1337–46. Smith, R. M., Chienthavorn, O., Wilson, I. D., Wright, B. and Taylor, S. D. (1999), ‘Superheated heavy water as the eluent for HPLC – NMR and HPLC – NMR – MS of model drugs’, Anal. Chem., 71, 4493–97. Smith, R. M. (2002), ‘Extractions with superheated water’, J. Chromatogr. A, 975, 31–46. Sporring, S. and Björklund, E. (2004), ‘Selective pressurized liquid extraction of polychlorinated biphenyls from fat-containing food and feed samples. Influence of cell dimensions, solvent type, temperature and flush volume’, J. Chromatogr. A, 1040, 155–61. Tajuddin, R. and Smith, R. M. (2002), ‘On-line coupled superheated water extraction (SWE) and superheated water chromatography (SWC)’, Analyst, 127, 883–85. Tanaka, T., Hori, T., Asada, T., Oikawa, K. and Kawata, K. (2007), ‘Simple one-step extraction and cleanup by pressurized liquid extraction for gas chromatographic-mass spectrometric determination of pesticides in green leafy vegetables’, J. Chromatogr. A, 1175, 181–86. Turner, C. and Mathiasson, L. (2000), ‘Determination of vitamins A and E in milk powder using supercritical fluid extraction for sample clean-up’, J. Chromatogr. A, 874, 275–83. Turner, C., Turner, P., Jacobson, G., Almgren, K., Waldeback, M., Sjöberg, P., Nordberg-Karlsson, E. and Markides, K. E. (2006), ‘Subcritical water extraction and b-glucosidase-catalyzed hydrolysis of quercetin glycosides in onion waste’, Green Chem., 8, 949–59. Vandenburg, H. J., Clifford, A. A., Bartle, K. D., Zhu, S. A., Carroll, J., Newton, I. D. and Garden, L. M. (1998), ‘Factors affecting high-pressure solvent extraction (accelerated solvent extraction) of additives from polymers’, Anal. Chem., 70, 1943–48. Vandenburg, H. J., Clifford, A. A., Bartle, K. D., Carroll, J. and Newton, I. D. (1999), ‘Comparison of pressurised fluid extraction and microwave assisted extraction with atmospheric pressure methods for extraction of additives from polypropylene’, Analyst, 124, 397–400. Wahlen, R. and Catterick, T. (2004), ‘Simultaneous co-extraction of organometallic species of different elements by accelerated solvent extraction and analysis by inductively coupled plasma mass spectrometry coupled to liquid and gas chromatography’, Rapid Commun. Mass Spectrom., 18, 211–17. © Woodhead Publishing Limited, 2010
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Principles of physically assisted extractions and applications 71
3 Principles of physically assisted extractions and applications in the food, beverage and nutraceutical industries E. Vorobiev, Compiègne University of Technology, France and F. Chemat, University of Avignon and Pays de Vaucluse, France Abstract: The physical methods used to enhance pressure and solute extraction methods in the food industry are reviewed. The quality of extracted products (e.g. purity, colour, texture, flavour and nutrients) may be degraded by conventional mechanical, thermal or chemical pretreatments. Several physical treatments (pulsed electric fields, power ultrasound, microwaves, ohmic heating, arc discharges) have been shown to be important and of great interest for the food industry in terms of enhancing solute and pressure extraction, and dehydration processes. Emerging extraction technologies that show promise for commercial food processing are discussed. Key words: plants, food, extraction, pulsed electric field, ultrasonic extraction, microwave extraction.
3.1 Introduction Pressure and solvent extractions are widely used in the food industry for the production of juices, wine, sugar and vegetable oil; these methods are also frequently applied in the extraction of various target compounds, such as colorants, antioxidants, essential oils, and aromas, from raw plant materials. Pressing, hot water and organic solvent extractions are long established processes that have excellent efficacy when applied in optimal conditions. These extractions can be done concurrently but more often they compliment each other and are technologically combined (e.g. the hot water extraction of sugar from sugar beets is combined with the subsequent pressing of pulps; the pressing of oilseeds is combined with the subsequent solvent extraction of oil from press-cake; the pressing of apples or grapes can be combined © Woodhead Publishing Limited, 2010
72 Separation, extraction and concentration processes with the subsequent solvent extraction of bioactive compounds from mash). The yield of extracted compounds can be very high in optimal conditions. Unfortunately, the quality of solutions and extracted products (e.g. purity, turbidity, colour, texture flavour and nutrients) may be degraded in the course of the raw material treatments that are necessary to increase the yield (grinding, heating, chemicals/enzyme addition). Moreover, a significant quantity of waste is often produced from the purification of impure extraction solutions, when undesirable loss of solvents and other additives occurs. These problems need to be resolved or minimised for the future development of environmentally sustainable food technologies. In recent decades there has been a growing interest in alternative food technologies that allow for nonthermal or mild thermal food preservation. Several emerging technologies are prominent and of great interest to the food industry, in particular pulsed electric fields (PEF), power ultrasound (PU), microwave (MW), pulsed light (PL), ohmic heating (OH), irradiation (IR), radio-frequency (RF) heating, high pressure processing (HP), high voltage electric discharges (HVED), amongst others (Barbosa-Cánovas et al., 1998; Zeuthen and Bogh-Sorensen, 2000; Povey and Mason, 1998). Such success in the development of novel preservation technologies has encouraged further research and has renewed the industrial interest in extractions assisted by nonthermal or mild thermal physical treatments (PU, PEF, HVED, OH, MW). Recently, these treatments were found to be effective for the enhancement of solute and pressure extractions, and dehydration processes (Knorr et al., 2001; Kulshrestha and Sastry, 2003; Li et al., 2004; Virot et al., 2007; Vorobiev and Lebovka, 2006, 2008). The potential benefits of physically assisted extractions are important. For instance, they can lead to the future substitution of hot water or organic solvent extractions with cold, or mild thermal, aqueous or pressure extractions. Moreover, the alternative physical treatments are found to be less invasive methods for the processing of plant foods, making it possible to avoid many undesirable changes in products, pigments, vitamins, and flavouring agents, which are typical of other extraction techniques, including thermal, chemical and enzymatic techniques. Some of these novel technologies remain very much in the research arena, whereas some others are on the brink of commercialisation. This chapter provides an overview of the emerging extraction technologies that are promising in terms of their potential applications in commercial food processing.
3.2 Pulsed electric field-assisted extractions in the food industry 3.2.1 Principles of pulsed electric field (PEF) treatment Pulsed electric field processing is a technique in which a food is placed between two electrodes in a batch or continuous treatment chamber and © Woodhead Publishing Limited, 2010
Principles of physically assisted extractions and applications 73 exposed to a pulsed voltage (typically 15–80 kV cm–1 with several pulses of 1–5 ms for the killing of micro-organisms; and 0.1–5 kV cm–1 with pulses of 50–1000 ms for the electroporation of plant cells and nonthermal extraction from solid foods). To generate such short pulses, various pulse-forming networks are used, with main components that include power supply at the selected voltage, one or several capacitor banks, inductors and/or resistors (Barbosa-Cánovas et al., 1998; Bluhm, 2006). Various different pulse shapes can be generated, including the simplest exponential decay pulses and squarewave pulses (mono-and bipolar). The duration and/or number of pulses are limited in order to prevent any significant temperature elevation, which is usually less than 3–5 °C. Mechanism of cell membrane electroporation When the intensity of an applied electric field increases, the potential difference across a cell membrane (transmembrane potential um) also increases. If um exceeds a stated threshold value (typically 0.2–1 V), a temporary loss of membrane semipermeability occurs. This phenomenon of membrane damage is named electroporation (electropermeabilisation) (Fig. 3.1). A number of models have been proposed in the biophysical literature to describe the electroporation mechanism (Konduser and Miklavcic, 2008; Zimmermann,
E < Ecr
Compression
Compressed membrane
E > Ecr
(a) Repulsion
Pore
Membrane
2r
(b)
Fig. 3.1 Membrane electroporation. With application of electric field E, the cell membrane is compressed owing to the electric charges attraction (a). (b) When the electric field exceeds some critical value (E > Ecr), the transmembrane potential increases to a threshold value and electroporation occurs (from Vorobiev et al., 2005).
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74 Separation, extraction and concentration processes 1986). The transmembrane potential um of a spherical cell depends on the angle q between the external field E direction and the normal vector on the membrane surface. The transmembrane potential may be given by the following equation (Zimmermann, 1986):
um = 0.75fdcEcosq [1 – exp (– t/tc)]
[3.1]
where dc is the cell diameter, tc is the time of membrane charging, and f is a parameter depending on the electrophysical and dimensional properties of the cell, membrane, and surrounding media. The values of f and tc vary for different cell sizes and ratios of extracellular/intracellular conductivities se/ si (Lebovka et al., 2000; Vorobiev and Lebovka, 2008). At complete cell damage, it follows that se/si = 1 and f ≈ 1. The critical value of transmembrane potential required for biological membrane electroporation at ambient temperature is estimated as um ≈ 1 V (Zimmerman, 1986) and can vary in plant and animal cells from 0.7 to 2.2 V (Knorr et al., 2001). The typical size of a cell in plant tissue is dc ≈ 50–100 mm. Therefore, the required threshold PEF intensity for the electroporation of food plants tissues (for f = 1 and tc <
Z = (s – si)/(sd – si)
[3.2]
where s is the electrical conductivity value measured at low frequency (1–5 kHz) and subscripts ‘i’ and ‘d’ refer to the conductivities of intact and totally damaged cells, respectively. This equation gives Z = 0 for intact tissue and Z = 1 for totally disintegrated material. Another method is based on electrical conductivity measurements at low (ª1 kHz) and high (3–50 MHz) frequencies (Angersbach et al., 2002). Figure 3.2 shows the characteristic time tE that is necessary for halfdamage (i.e. Z = 0.5) of red beet tissue (Shynkaryk et al., 2008). Inserts show examples of the conductivity disintegration index Z as a function of
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Principles of physically assisted extractions and applications 75 Conductivity disintegration index, Z
101
100
Time, tE (s)
10–1
10–2
1
E = 600 V cm–1
400 V cm–1 300 V cm–1 200 V cm–1
0.5
tE
10–5
10–4 10–3
10–2 10–1 tt (s)
100
101
10–3
10–4
10–5
200
300 400 500 600 Electric field strength, E (V cm–1)
700
Fig. 3.2 Characteristic time of half-damage (Z = 0.5) of red beet tissue (Shynkaryk et al., 2008), with insert showing examples of the conductivity disintegration index Z as a function of PEF treatment time.
treatment time tt = nNti, where ti = 100 ms is the duration of the square pulse, n is the number of pulses, and N is the number of trains. The PEF treatment was conducted at a room temperature, T = 20 °C. The value of tE decreases with the increase of the electric field strength E and its approaching is tE ª10–4 s at E ≥ 600 V cm–1. The compression-to-failure and the stress–relaxation measurements for apple, carrot and potato tissues treated using PEF over different time periods tPEF were presented by Lebovka et al. (2004b). After undergoing PEF treatment of rather high intensity and long duration (E = 1.1 kV cm–1, tPEF = 0.1 s), the tissues had lost part of their initial strength. However, the changes of both the elasticity modulus Gm and the fracture stress PF were significantly smaller than the changes observed for the freeze-thawed and thermally (T = 45 °C, 2 h) pre-treated tissues. It was concluded that PEF enables a high disintegration of membranes to be achieved, and a turgor component to be removed from the texture. However, compared with tissue treated by freeze-thawing or heating, the tissue structure seems to be less affected by PEF treatment. This conclusion was later confirmed by the textural studies of sugar beet tissue treated by PEF (Shynkaryk et al., 2008). Microscopic studies of onion tissue treated by PEF have shown intact cell architecture, whereas colorant could easily diffuse inside of cells with electroporated membranes (Fincan and Dejmek, 2002). Lebovka et al. (2000, 2001) have put forward a hypothesis on how PEF
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76 Separation, extraction and concentration processes treatment affects cellular tissues. They consider the effect to be correlated percolation which is governed by two processes. The first of these is the resealing of cells, and the other is the transfer of moisture inside the cellular structure, which is sensitive to repetition of PEF treatments. When the treatment is applied using a low enough intensity of electric fields, the electroporation is reversible as the resealing process is quick to repair the membranes immediately after the PEF treatment has been terminated. At moderate PEF treatment, some of the cells lose their permeability, but in others the pores may persist (Lebovka et al., 2001). High-intensity PEF treatment causes irreversible damage to the cell membrane. Long-term changes in conductivity after the application of PEF treatment can also be related to osmotic flow and to the redistribution of moisture inside the sample (Lebovka et al., 2001). PEF-treatment chambers Special chambers (batch or continuous) have been developed for PEF-assisted extraction (Bluhm and Sack, 2008; Jaeger et al., 2008; Lazarenko et al., 1977; Vorobiev and Lebovka, 2008). The chambers specified for the treatment of liquid foods by high PEF (Barbosa-Cánovas et al., 1998) can also be used for the pretreatment of yeasts before extraction (Shynkaryk et al., 2008). The chambers developed for the sterilisation of particulate foods by ohmic heating (Biss et al., 1989) can be adapted for the PEF pretreatment applied before the extraction of apple mash (Jaeger et al., 2008). In some cases, the PEF treatment can be combined with extraction in the same apparatus (Vorobiev and Lebovka, 2008). Figure 3.3(a) and (b) presents the laboratory treatment cell developed in Compiegne University of Technology, which combines PEF treatment and juice expression from solid foods. 3.2.2 Juice expression and solute extraction from plant roots, tubers and fruits In raw food plants, the valuable compounds are initially enclosed in cells, which have to be damaged to facilitate the recovery of intracellular matter. Conventional techniques of cell damage, such as fine mechanical fragmentation, and thermal, chemical and enzymatic treatments, lead to the more severe disintegration of tissue components, including cell walls and cell membranes. PEF treatment, which is less destructive than conventional methods, can be used for the more selective extraction of cell components. Sugar beets The conventional extraction technology consists of a power-consuming hot water diffusion of sugar from sliced cossettes at 70–75 °C. Unfortunately, the denaturation of tissue by heat causes alteration in the cell wall structure through hydrolytic degradation reactions. In addition to sucrose, other cell components, such as pectin pass in juice during extraction, affecting the juice
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Principles of physically assisted extractions and applications 77 PEF generator
Data processing
Compressed air
Steel frame Elastic rubber Pressure
Polypropylene ring Mobile electrode Sample Stationary wire gauze electrodes
Rubber shims
Filter cloth
Steel cover Juice treatment cell (a)
(b)
Fig. 3.3 Laboratory cell combining PEF treatment and juice expression from solid foods: (a) schematic diagram; (b) photographs.
purity. Consequently, a complex multi-staged juice purification process is used in beet sugar production (Van der Poel et al., 1998). The PEF application has great potential as an alternative method to the conventional thermal
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78 Separation, extraction and concentration processes technology. Recently, PEF-assisted pressing and aqueous extraction from sugar beets have been intensively studied (Bouzrara and Vorobiev, 2000; Bouzrara, 2001; El-Belghiti and Vorobiev, 2004, 2005a; Eshtiaghi and Knorr, 2002; Jemai and Vorobiev, 2003, 2006; Lebovka et al., 2007). Several studies have demonstrated the efficiency of PEF treatment for the cold pressing of sugar beet cossettes (Bouzrara and Vorobiev 2000; Bouzrara, 2001; Eshtiaghi and Knorr, 2002). Bouzrara and Vorobiev (2000) and Bouzrara (2001) thus reported that up to around 82% of the overall yield could be achieved by two-stage pressing with intermediately applied PEF (E = 400 V cm–1, tCEP = 0.1 s). The initial pressurisation of slices serves to assure a good electrical contact between them. In addition, the secondary juices (i.e. after PEF application) were systematically more concentrated in sugar and had lower colour concentration. PEF treatment was successfully applied with scale-up experiments exploring the viability of a novel process of cold juice extraction (Jemai and Vorobiev 2006). The processing scheme consisted of two initial pressing steps with an intermediate PEF treatment, followed by one or more washing steps and a final pulp pressing. The cold juices expressed from sugar beet gratings after the intermediate PEF treatment have higher purity values (95~98%) when compared with those before PEF application (90~93%). Additionally, the quantity of pectin was noticeably lower and the colour concentration of the juice was systematically 3 to 4 times lower than the colour of factory juices (Jemai and Vorobiev, 2006). These results, which show significant amelioration of the qualitative juice characteristics, give interesting new perspectives on cold PEF-enhanced expression from the sugar beets. Figure 3.4 shows a schematic diagram and a photograph of a pilot belt press that has recently been used for the PEF-assisted expression from sugar beets (Grimi et al., 2008). The results presented in Fig. 3.4 confirm the amelioration of juice yield (Fig. 3.4c) and purity (Fig. 3.4d) with the application of PEF pre-treatment. The size of particles treated by PEF should be optimised to obtain the juice with a maximal yield and higher purity level. Alternative studies have demonstrated the possibility of sugar extraction in cold or moderately heated water (Jemai and Vorobiev, 2003). Lebovka et al. (2007) compared the kinetics of thermal and cold diffusion and diffusion coefficients Deff for untreated and PEF-treated (E = 400 V cm–1, tCEP = 0.1 s) sugar beet slices (1.5 mm ¥ 10 mm ¥ 10 mm). The difference in Deff values between the untreated and PEF-pretreated tissue increases significantly for less heated tissue. For instance, the values of Deff were nearly the same for sugar diffusion at 60 °C from untreated tissue and at 30 °C from PEFpretreated tissue. The purest juice was obtained after cold diffusion. However, even after thermal diffusion at 70 °C, the juice purity was higher in slices pretreated by PEF than for untreated slices. Encouraging results obtained by several research groups revealed strong industrial interest, and semi-industrial scale equipment was built for PEFassisted extraction from sugar beets (Bluhm and Sack, 2008).
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Principles of physically assisted extractions and applications 79 Other roots and tubers In red beetroot, the water-soluble betalaines (red-violet betacyanins and yellow betaxanthins) are the main red–purple pigments. The high degree of extractability from red beetroot was observed after the PEF treatment at field strength 1 kV cm–1, when samples released about 90% of the total red
Upper belt F
C
J Drainage area
Low belt
J (a)
(b)
Fig. 3.4 Pilot belt press: (a) schematic diagram (F, feeding; C, cake; J, juice extract); (b) photograph; (c) juice yield; and (d) juice purity (from untreated and PEF-treated sugar beet slices of different sizes: S1 0.045 mm3, S2 47.5 mm3, S3 280 mm3 and S4 1050 mm3).
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80 Separation, extraction and concentration processes 100 Untreated PEF treated
Juice yield (%)
80
60
40
20 0
S 1
<
S 2
< (c)
S 3
<
S4
100
Purity (%)
Untreated PEF treated
95 90 85 80
S 1
<
S 2
< (d)
S 3
<
S4
Fig. 3.4 Continued
pigment following 1 h of aqueous extraction (Fincan et al., 2004). Lopez et al. (2009a) have shown that the application of PEF treatment at 7 kV cm–1 facilitates an increase in the maximum yield of betanine by a factor of 4.2, compared with PEF treated samples, achieving almost complete betanine release. The combination of PEF at 7 kV cm–1 and pressing at 14 kg cm–2 shortened the extraction time 18-fold. It can be noted that the effective electroporation of red beet tissue at ambient temperature can be attained even at the lower electric fields strengths of 400–600 V cm–1 (Shynkaryk et al., 2008; Fig. 3.2). For carrots, the PEF-assisted expression (Bouzrara, 2001; Knorr et al., 1994; Praporscic et al., 2007b), aqueous extraction (El-Belghiti and Vorobiev, 2005b) and their combination (Grimi et al., 2007) were studied. A significant enhancement of juice yield can be attained, even at a rather low voltage gradient of 360 V cm–1 (Bouzrara, 2001). Juices expressed with PEF treatment were more transparent and less turbid than untreated juices. Moreover, the Brix values were instantly increased after the PEF application (Praporscic et al., 2007b). Grimi et al. (2007) studied carrot juice extraction using a laboratory filter-press chamber with various combinations of pressing © Woodhead Publishing Limited, 2010
Principles of physically assisted extractions and applications 81 and washing operations. They have shown that it is possible to produce from the press-cake a ‘sugar-free’ concentrate, rich in vitamins and carotenoids, that can be used as an additive in diet foods. Potato tissue can also be effectively electroporated by PEF. Its textural and compressive properties after PEF treatment were studied by Lebovka et al. (2004b, 2007), and Grimi et al. (2009a). A prototype for potato starch extraction by PEF was developed by ProPuls, Germany, and commercial prototypes for this application based on a patented Marx Generator design have been developed by KEA-Tec, Germany (Jaeger et al., 2008). Apples Various research groups have studied the influence of PEF treatment on the expression of apple juice (Bazhal and Vorobiev, 2000; Bazhal et al., 2001; Lebovka et al., 2004a; McLellan et al., 1991; Praporscic et al., 2007b; Schilling et al., 2007). The reported improvement of juice yield varied, probably owing to the different processing conditions employed in the different studies (degree of particle fragmentation, PEF parameters, and compression pressure). For instance, the size of particles and fragmentation method (slicing, grinding or milling) can be essential for the improvement of juice yield. Recently, Grimi et al. (2008) demonstrated that the juice yield obtained from Golden Delicious apple slices (2 ¥ 3.5 ¥ 55 mm) after PEF treatment (E = 400 V cm–1, tCEP = 0.1 s) increased by 28%, whereas it increased by just 5% in finer slices (1 ¥ 1.9 ¥ 55 mm). The finely fragmented particles cause most cells to be disrupted mechanically, and the additional effect of the PEF on the total juice yield is rather limited. On the contrary, with coarse particles, the percentage of cell membranes damaged electrically increases, but the cell wall structure is less affected. This might be the reason for the more transparent and less cloudy apple juices obtained after PEF treatment of coarse particles (Bazhal and Vorobiev, 2000; Praporscic et al., 2007b). No apparent change in pH value or total acidity was detected. Additionally, the content of many nutritionally valuable compounds was retained, or even enhanced (Jaeger et al., 2008).The good industrial potential for PEF-assisted apple juice expression is confirmed in the laboratory and pilot scales using a belt-press equipment (Grimi et al., 2008; Jaeger et al., 2008). In addition, Jemai and Vorobiev (2002) have demonstrated that PEF exerts an enhancing effect on the diffusion of soluble substances from apple tissue. Apple discs (Golden Delicious) were pretreated thermally (75 °C, 2 min) and by PEF (E = 100–500 V cm–1, tPEF = 0.1 s), and afterwards the diffusion kinetics was studied at various temperatures. Experiments revealed that a detectable enhancement of diffusion kinetics starts at field intensities of 100–150 V cm–1. A further increase of both the field intensity and the pulse duration led to an enhancement of the diffusion kinetics. Jemai and Vorobiev (2002) showed that for thermally treated samples, the temperature variation of the diffusion coefficient D is of Arrhenius type with two diffusion
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82 Separation, extraction and concentration processes regimes: (i) without thermal pre-treatment (Ea~28 kJ mole–1) and (ii) after thermal denaturation (Ea~13 kJ mole–1). Only one regime with intermediate activation energy (Ea~20 kJ mole–1) was observed for electrically treated samples. Grapes Electroporation of wine grapes is an alternative nonthermal process leading to a prudent extraction of colours and valuable constituents. Recently, Praporscic et al. (2007a) investigated quantitative (juice yield) and qualitative (absorbance and turbidity) characteristics of juices obtained during the expression of white grapes (Muscadelle, Sauvignon and Semillon). The experiments were carried out at an expression pressure of 5 bar using a laboratory compression chamber equipped with a PEF treatment system. PEF treatment with a field strength of E = 750 V cm–1 and the total treatment duration tPEF = 0.3 s was applied. The PEF treatment resulted in an increase of the final juice yield Yf of up to 73–78% compared with Yf ª 49–54% for the untreated grapes. A rather noticeable decrease of absorbance and turbidity was observed as a result of the PEF treatment for all the white grape varieties studied. Furthermore, Grimi et al. (2009b) have shown that PEF treatment enhances compression kinetics and extraction of polyphenols from Chardonnay grapes. Lopez et al. (2009b) have studied the application of a PEF treatment (5 kV cm–1, 50 pulses) to the de-stemmed, crushed and slightly compressed grape pomace (skins, pulp and seeds) of Cabernet Sauvignon grapes. The application of a PEF treatment to the pomace before the vinification process led to freshly fermented wines that were richer in colour intensity, contained more anthocyanins and tannins, and showed better visual characteristics. The PEF treatment has facilitated a reduction in maceration time during the vinification of Cabernet Sauvignon grapes from 268 to 72 h. A specially designed mobile facility was built to produce electric fields of up to 60 kV cm–1 in the reactor at a repetition rate of up to 15 Hz. It is able to handle a throughput of up to 1 t h–1 with an energy consumption of around 15 kWh t–1. Both red and white wine grapes have been treated (Bluhm and Sack, 2008). Yeast cells The disruption of yeast cells (Saccharomyces cerevisiae) is a very important step in the industrial extraction of bioproducts, such as valuable proteins, cytoplasmic enzymes, and polysaccharides), which are present inside the cells. A thermal treatment at T > 50 °C results in damage to the yeast membranes, but also causes the denaturation and degradation of many valuable intracellular components. Electroporation exerts a minimal undesirable impact on liquid components inside and outside the cell, and can be done without significant temperature increase and cell debris formation. The commonly reported values of field strength E needed for the disintegration of membranes in S. cerevisiae cells, are rather high (E > 7.5 kV cm–1) (Schrive et al., 2006;
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Principles of physically assisted extractions and applications 83 Zhang et al., 1994), though smaller fields can also affect the structure of cells. For instance, the early stages of damage in the yeast cells were observed at E < 7.5 kV cm–1, in a PEF treatment of long duration (El Zakhem et al., 2006a,b). The electroextraction of proteins from suspensions of S. cerevisiae was also observed at 3.2 kV cm–1 (Ganeva et al., 2003), but high efficiency (yield of 85%) required long extraction following PEF treatment (>4 h at 30 °C). Recently, Shynkaryk et al. (2009) demonstrated that a rather long application of PEF treatment (1000 trains, 200 pulses of 100 ms in each train) was needed to result in a high level of membrane disintegration (Z > 0.8), even at E = 10 kV cm–1. Moreover, the quantity of released high molecular weight intracellular components extracted with PEF was lower than that extracted with other methods (HVED and high-pressure homogenisation) (Shynkaryk et al., 2009).
3.3 Ohmic heating-assisted extractions in the food industry 3.3.1 Principles of ohmic heating (OH) treatment Ohmic heating (OH) processing is a technique in which a food is placed between two electrodes in a batch or continuous treatment chamber and exposed to a dc, ac or pulsed voltage (typically 20–80 V cm–1) to heat the food and kill any micro-organisms. OH generators, which use ac and highfrequency electric fields permit reduced electrolysis and product contamination, compared with dc designs (Lima et al., 1999; Sastry, 2005). The basic principle of OH is the dissipation of electrical energy in the form of heat, resulting in the generation of internal energy (Sastry, 2005):
q = sE2
[3.3]
where q is the heat power per unit volume, s is the electrical conductivity of food and E is the electric field intensity. OH processing is currently commercialised for the pasteurisation and sterilisation of particulate foods (Biss et al., 1989; Sastry, 2005). However, the effects of OH are not purely thermal. Several studies have noted that electropermeabilisation is also induced by OH (Kemp and Fryer, 2007; Kulshrestha and Sastry, 2003; Lima et al., 2001; Praporscic et al., 2006; Schreier et al., 1993). Lebovka et al. (2005) have studied the membrane damage of ohmically heated potato and apple tissues. The OH experiments conducted under ac treatment were started from room temperature T = 22 °C at different electric field strengths E, and the treatment was stopped when the temperature reached T = Tf = 50 °C. The total treatment time te was higher with lower E (Fig. 3.5), but the samples were not kept in the warm juice (at 40–50 °C) for more than 15 min. The thermal damage and softening of potato and apple tissues at such mild thermal treatment conditions can be neglected (Lebovka et al., 2004a; 2004b). But the conductivity disintegration index Z © Woodhead Publishing Limited, 2010
84 Separation, extraction and concentration processes increased considerably at the electric fields E > 20 V cm–1, both for apples and potatoes (Fig. 3.5), and these data provide evidence for the significance of electrically induced tissue damage. More recently, Lebovka et al. (2007) estimated the variation of Z values of sugar beet tissue during OH. They obtained electrical conductivity graphs for intact, maximally damaged and ohmically treated tissues at various temperatures and calculated Z(T) from equation [3.2] for various electric fields E. The results provide evidence for the significance of the electroporation effect in OH. This effect becomes more significant as E is increased from 40 to 100 V cm–1. Imai et al. (1995) have found that the OH behaviour and impedance values of Japanese white radish depended on the frequency of the electric field, which varied from 50 Hz to 10 kHz. At a constant voltage of 40 V cm–1, the time required to heat the sample to the desired temperature was longer for higher frequencies of electric field. It was concluded that the initial rapid heating of white radish at low frequency (50 Hz) is caused by the electroporation effect. 3.3.2 Juice expression and solutes extraction enhanced by OH The first studies of extractions enhanced by OH were completed in the former Soviet Union, and summarised by Rogov and Gorbatov (1974) and by Lazarenko et al. (1977). In Soviet publications, electroporation was known as ‘electroplasmolysis’, and the combined thermal and electric effects caused by OH were called ‘thermo-electroplasmolysis’ or ‘low-gradient ac, Tf = 50 °C
Potato Apple 103
0.6
te (s)
Conductivity disintegration index, Z
0.8
0.4 102
0.2
0
0
20
40 60 80 Electric field strength, E (V cm–1)
100
101
Fig. 3.5 The conductivity disintegration index Z and time of treatment te versus electric field strength E for potato and apple tissues ohmically heated to 50 °C by ac treatment.
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Principles of physically assisted extractions and applications 85 plasmolysis’ to underline the relatively low electric fields (typically 10–100 V cm–1) used for OH. Research interest in extractions assisted by OH processing was renewed in Europe and the USA during the 1990s and 2000s (Sastry, 2005; Schreier et al., 1993; Wang and Sastry, 2002) probably due to the successful commercialisation of OH systems for food sterilisation (Biss et al., 1989). This interest in OH extractions focused on the synergy of electrical and thermal effects on cell tissue, and the lower temperatures needed for effective membrane damage compared with conventional heating. On the other hand, the electrical field applied in OH-assisted extraction is lower than that for nonthermal PEF treatment. Therefore, electrical equipment can be simplified to deliver a lower output voltage than typical high voltage PEF generators. Various studies have recently demonstrated the effectiveness of OH pretreatment for the intensification of pressing and solute extraction (Kemp and Fryer, 2007; Kulshrestha and Sastry, 2003; Lima et al., 1999; Praporscic et al., 2005, 2006; Sensoy and Sastry, 2004; Wang and Sastry, 2002). For instance, Lima et al. (1999) ohmically heated apple cuts up to 40 °C with a voltage of E = 40 V cm–1 and wave frequencies of 4 and 60 Hz. Subsequent pressing at a constant rate (3 cm min–1) resulted in a higher juice yield than that obtained for untreated cuts, and the 4 Hz pretreatment was more effective than the 60 Hz pretreatment. More recently, Wang and Sastry (2002) have demonstrated that OH pretreatments of 40 or 50 °C by ac with 40 V cm–1 and 60 Hz enhanced the juice yield from apple samples. Compared with untreated samples, or those preheated by microwaves, the juices extracted from ohmically preheated apple tissue were visually similar in quality, and the yields were higher. The lower frequency of OH compared with MW heating may be the important factor that caused the greater increase in juice yield (Wang and Sastry, 2002). There may be other mechanisms involved in intra- and extracellular moisture flow during OH. In addition to its thermal effects and electrical breakdown, electro-osmosis may also be the reason for moisture flowing inside of cell tissue. Bazhal and Vorobiev (2000) have demonstrated that electro-osmotic flow was induced in apple tissue by a dc electric field of low voltage, which resulted in a significant increase in juice yield, even at ambient temperature. Solute extraction can also be enhanced by OH pretreatment. Kemp and Fryer (2007), Kulshrestha and Sastry (2003), and Schreier et al. (1993), showed that there is an increased diffusion of betanine dye from red beets treated by OH. In their interesting study, Kulshrestha and Sastry (2003) treated red beet cuts at 45 °C over a range of voltages from 0 to 23.9 V cm–1, and at frequencies of 0 (dc), 10, 50, 250 and 5000 Hz. As the electric field frequency decreased, the betanine content in the solution increased. This was explained by the increased time for which the cell membrane was charged at lower field frequencies. An exception to this tendency was detected for those tissues treated by dc, which was less effective than OH treatment at 10 Hz. All the mechanisms involved in tissue modification by
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86 Separation, extraction and concentration processes OH treatment have not yet been elucidated. Electrolysis phenomena might also be important in terms of process optimisation. Other fresh materials, for example, mint leaf (Sensoy and Sastry, 2004), potato (Praporscic et al., 2006), and sugar beets (Praporscic et al., 2005) can also be processed by OH-assisted extraction.
3.4 Extraction assisted by high-voltage electrical discharges and applications in the food industry 3.4.1 Principles of underwater high-voltage electrical discharges (HVED) Underwater (electrohydraulic) HVED is a technique in which solid particles are placed in a dielectric liquid (typically tap water) inside a chamber, containing a high-voltage (HV) needle electrode and a plated grounded electrode, and exposed to pulsed shockwaves (typically 40–60 kV cm–1, 2–5 ms) to produce a liquid breakdown and particle fragmentation. This technique has various applications (defragmentation of solid materials, such as concrete, treatment of municipal solid waste, and drilling) (Bluhm, 2006). Some publications propose the use of the HVED application for the deactivation of micro-organisms and for the enhancement of extractions (Grémy-Gros et al., 2008). Figure 3.6 shows a schematic diagram of a HVED generator, as well as typical voltage and current curves measured during a HVED (Boussetta et al., 2009a). Effects of HVED The mechanisms of HVED are very complex and not yet well understood. The phenomenon is based on the electrical breakdown of water (Bluhm, 2006). Air bubbles that are already present in the water, or formed as a result of local heating, participate in and accelerate this phenomenon. If the electrical field is intense enough, the avalanche of electrons becomes a starting point for streamer propagation from the HV needle electrode to the grounded one (Fig. 3.7). A crack can be attracted from the discharge channel into the solid inclusion if their dielectric properties are different. The electrical breakdown is accompanied by a number of secondary phenomena (high-amplitude pressure shock waves, bubbles cavitation, and creation of liquid turbulence). For instance, the pressure exerted by the expanding channel almost always exceeds the tensile strength of the solid inclusions, and leads to the formation of cracks in the solids. Several hundred to thousands of discharges are sufficient to fragment the stones into small pieces of less than 1 mm (Bluhm, 2006). For biological particles, the cell structure can be damaged by HVED, which accelerates solute extraction (Grémy-Gros et al., 2008). If electrical breakdown does not occur, water carries electricity and behaves
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Principles of physically assisted extractions and applications 87 HVED generator
Vent hole
Insulation
Air Water + sample
Needle electrode Plate electrode
Metal
HVED treatment cell (a) 15
40
10
20 5
0
0
Current (kA)
Voltage (kV)
tf ª 0.5 ms
–5 0
2
4 6 Time (ms) (b)
8
10
Fig. 3.6 HVED treatment (a) experimental setup and (b) typical voltage and current graphs.
like a good electrical conductor; then electrical signals have exponential form. When HVED are produced, electrical signals look like damped oscillations, as presented in Fig. 3.6 (Boussetta et al., 2009a; Grémy-Gros et al., 2008). 3.4.2 Solute extraction enhanced by HVED HVED have been used to accelerate the extraction of solutes from soybeans, potato, tea leaves, peat, and fennel (Barskaya et al., 2000; El-Belghiti, 2005; El-Belghiti et al., 2007) as well as from other food materials.
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88 Separation, extraction and concentration processes
(a)
(b)
(c)
HV electrode
Shock wave
Charging unit
Ground Solid (d)
Fig. 3.7 Type of electric discharge in aqueous solution: (a) streamer, E = 4.4 kV cm–1; (b) streamer and spark, E = 13.3 kV cm–1; (c) spark, E = 33.3 kV cm–1 (Sugiarto et al., 2001). (d) Schematic diagram of electro-hydraulic disintegration process.
Oil The application of HVED can enhance the aqueous extraction of oil from oilseeds (Grémy-Gros et al., 2008; Gros et al., 2003). Linseeds were crushed and pressed with a hydraulic press, and then the press-cake obtained was reduced to powder and dissolved in demineralised water at ambient temperature. The mixture was treated with HVED (1–1640 discharges) and centrifuged (Gros et al., 2003). Approximately 26% of oil remained in the residue after 1640 pulses. The process was then optimised by varying the number of pulses, pH, water/press-cake ratio and temperature (Grémy-Gros et al., 2008). The HVED treatment also made it possible to enhance mucilage extraction from whole linseed (Gros et al., 2003). Seeds were dissolved in demineralised water and treated for 10 min (i.e. 300 pulses of 0.5 Hz frequency). A centrifuge separation was then performed to obtain solution and residue. The residue was then treated a second and a third time with fresh water under the same conditions. During the HVED treatment, linseeds were defragmented under the action of shock waves. Three successive 10-min treatments were sufficient to extract the mucilage almost entirely. The extraction of proteins began during the third treatment. However, the oil–protein–water emulsions obtained by HVED treatment then had to be separated. The following research study
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Principles of physically assisted extractions and applications 89 resulted in the development of membrane separation of emulsions created by HVED (Li et al., 2009). Grape pomace Grape pomace (composed of stems, seeds and skins) has a high polyphenol content. These polyphenols have attracted great interest as they exhibit antibacterial, antiviral, and antioxidant properties, and can help prevent cardiovascular diseases (Dugand, 1980). Boussetta et al. (2009b) studied the application of HVED in the extraction of polyphenols from the pomace obtained as a residue from pressed white grapes (Vitis vinifera L., cultivar ‘Chardonnay’, vintage 2007). The treatment chamber (Fig. 3.6) was initially filled with pomace mixed with distilled water (the liquid-to-pomace ratio was 3:1, the water temperature was 20, 40 or 60 °C), and 80 successive discharges were applied (40 kV, pulse repetition rate 0.5 Hz). Afterwards, the total solute and polyphenol diffusion from the untreated and treated pomace mixtures were studied at the desired temperatures. Figure 3.8 shows that both temperature and HVED treatment improved the extraction of polyphenols. The difference between the yield of solutes for experiments with HVED treatment and without HVED treatment decreased with temperature elevation (Boussetta et al., 2009b). On the contrary, the opposite tendency was observed for the yield of polyphenols. The application of HVED enhanced the polyphenol yield, not just for fresh grape pomace, but also for sulfured
1
20
T (°C) 40
60 With HVED treatment
0.8
Ypolyphenols (%)
0.6
0.4
Without HVED treatment 0.2
0.0034
0.0032 1/T (K)
0.003
Fig. 3.8 Effect of treatment temperature on final yield of polyphenols after 1 h of extraction of fresh grape pomace with and without HVED treatment.
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90 Separation, extraction and concentration processes and frozen pomaces. The final yields of solutes, reached after HVED and subsequent diffusion for 40 min were more than twice those reached after 240 min without HVED. Yeast cells Shynkaryk et al. (2008) studied the efficiency of HVED-assisted extraction from wine yeast cells (S. cerevisiae) in aqueous suspension (1% w/w). HVED induced damage in yeast membranes, leakage in cytoplasmic ions, and the release of intracellular bioproducts; it also resulted in an increase in the disintegration index Z. For HVED treatment at 40 kV, the high level of membrane disintegration (Z > 0.8) required more than 100 discharges, which corresponded to the effective time of HVED treatment between 100 and 200 ms. HVED treatment decreased the average cell size somewhat, owing to cell disruption. However, the quantity of high molecular weight intracellular components released, controlled by the absorbance analysis, was lower than that obtained by high-pressure homogenisation (Shynkaryk et al., 2008).
3.5 Ultrasound-assisted extraction (UAE) in the food industry 3.5.1 Principles of UAE Power ultrasound, which has frequencies between 20 kHz and 1 MHz, is now well known to have significant effects on the rate of various physical and chemical processes. Cleaning and solubilisation are its more developed applications, and a large variety of ultrasound baths exist for chemical laboratory use. The effect of ultrasonic waves on solid samples is widely used for the extraction of aromas from plant materials, or metal impurities from soils. Degassing and stripping are widely used for flavour analysis, and in environmental and polymer research. Other interesting ultrasound applications include homogenisation, emulsification, sieving, filtration, and crystallisation. The most interesting effect of ultrasound-based operational units is the reduction of processing time and the increase in product quality. All these effects are attributed to acoustic cavitation: when a liquid is irradiated by ultrasound, micro-bubbles form, grow and oscillate extremely fast, and eventually collapse powerfully (if the acoustic pressure is high enough). When the size of these bubbles reaches a critical point, they collapse during a compression cycle and release large amounts of energy. The temperature and pressure at the moment of collapse have been estimated to be up to 5000 K and 5000 atmospheres. This creates hotspots that are able to dramatically accelerate chemical reactivity in the medium. When these bubbles collapse onto the surface of a solid material, the high pressure and
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Principles of physically assisted extractions and applications 91 temperature released generate microjets directed towards the solid surface. These microjets are responsible for the degreasing effect of ultrasounds on metallic surfaces, which is widely used for cleaning materials. Another application of microjets in the food industry is the extraction of vegetal compounds (Mason, 1990; Suslick, 1988). As shown in Fig. 3.9, a cavitation bubble can be generated close to the surface of the plant material. Then, during a compression cycle, this bubble collapses and creates a microjet (400 km h–1) directed toward the plant matrix. The high pressure and temperature involved in this process will destroy the cell walls of the plant matrix, and its content can be released into the medium. Power ultrasound involves the mechanical and chemical effects of cavitation. The mechanism can be explained by two competing theories. The hot spot theory assumes that the high pressures and temperatures generated in the bubbles during the last nearly adiabatic compression, just before collapse, are responsible for the breakage of molecular bonds and the formation of radicals. On the other hand, the electrical theory describes micro-discharges, produced by the high electrical fields generated by the deformation and fragmentation of the bubbles. The most commonly used frequencies in sono-extraction are between 20 and 40 kHz. With higher frequencies, cavitation would be more difficult to induce because the cavitation bubbles need a slight delay to be initiated during the rarefaction cycle. The higher the frequency, the shorter the rarefaction cycle, so the less time available for the bubble to be created (Povey and Mason, 1998).
Cavitation bubble with negative pressure
Maximum and critical bubble size 2
1
A new cavitation bubble appears and cycle is repeated
3
4
Implosion of bubble by compression
Fig. 3.9 Ultrasonic cavitation phenomenon and its action on extraction of natural products.
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92 Separation, extraction and concentration processes 3.5.2 UAE reactors The two most common types of ultrasound equipment that are used for extraction are the ultrasonic cleaning bath and the more powerful probe system. For small extraction volumes, an ultrasound horn with its tip submerged in the fluid can be sufficient. Large volumes of fluids have to be sonicated in an ultrasound bath or in continuous or recycled-flow sono-reactors. There are also ultrasonic reactors conceived especially for solvent extraction on a laboratory scale (3 l), in a pilot plant (30 l) and on an industrial scale (150–1000 l). Pump systems are coupled to the ultrasonic reactor to stir the mixture and to empty the system at the end of the experiment. The intensity of the ultrasound is about 1 W cm–2, with a frequency of 25 kHz. In order to maintain constant temperature, the reactor is made of a double mantle into which cooling water can circulate. The main advantage of this type of apparatus is that the natural products and extraction solvent are mixed into a container, and the ultrasound is directly applied to the mixture (Fig. 3.10). Recently, the new methodology of continuous-flow systems has been used in analytical chemistry. Most UAE applications have been developed in discrete systems, using a bath or an ultrasonic probe, particularly in the extraction of food samples. Less frequent has been the design of on-line UAE systems in the same field. However, it is noted that the latter approach is considerably faster. It consists of an open system, in which fresh solvent flows continuously through the sample. This induces the displacement of mass transfer equilibria toward the solubilisation of analytes into the liquid phase (Priego-Capote and Luque de Castro, 2007). 3.5.3 Factors affecting UAE All the research carried out over the last 50 years strongly supports the importance of each identified factor in the ultrasound process, where frequency, intensity, treatment time, temperature, pressure, and treatment
(a)
(b)
Fig. 3.10 (a) Laboratory (3 l) and (b) industrial (150 l) ultrasound extraction reactors (with permission, www.etsreus.com).
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Principles of physically assisted extractions and applications 93 media define the extraction and the separation kinetics of ingredients from natural products (De Gennaro et al., 1999). Generally, the highest efficiency of UAE, in terms of the yield and composition of the extracts, can be achieved by increasing the ultrasound power, reducing the moisture of food matrices to enhance solvent–solid contact, and optimising the temperature to allow a shorter extraction time. Solvent choice is dictated by the solubility of the ingredients of interest, the interactions between the solvent and matrix, and the intensity of ultrasound cavitation phenomena in the solvent. The impact of ultrasound is attributed to intracellular cavitation, that is, micro-mechanical shocks that disrupt cellular structural and functional components up to the point of cell lysis (Butz and Tauscher, 2002). Figure 3.11 shows a proposed mechanism for ultrasound-induced cell damage. Cavitation is the formation, growth and, sometimes, the implosion of microbubbles created in a liquid when ultrasound waves propagate through it. The collapse of the bubbles leads to energy accumulation in hot spots, where temperatures of above 5000 °C and pressures of approximately 500 MPa have been measured. This phenomenon can cause enzyme inactivation through three mechanisms, which can act alone or combined. The first one is purely thermal, owing to the enormous temperatures achieved during cavitation. The second is caused by free radicals, generated by water sonolysis; and the third is the result of the mechanical forces (shear forces) created by microstreaming and shock waves (Raso et al., 1999). 3.5.4 UAE: main applications Essential oils and aromas UAE has also been developed for the extraction of essential oils from aromatic plants, such as peppermint leaves (Shotipruk et al., 2001), artemisia (Asfaw et al., 2005) and lavender (Da Porto et al., 2009), or from other vegetal matrices, such as garlic (Kimbaris et al., 2006) and citrus flowers (Alissandrakis et al., 2003). Increased yields of essential oil were found for peppermint leaves (up to 12%) and for artemisia when using UAE; and yields of the main compounds of lavandula essential oil were found to increase by Ultrasound (US)
US
H 2O
Cell
US
Cell
Pore initiation
US
H2O influx and swelling
Cell
Membrane rupture and cell lysis
Fig. 3.11 Mechanism of ultrasound-induced cell damage.
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94 Separation, extraction and concentration processes 2- to 3-fold when comparing UAE with conventional distillation. Moreover, UAE not only improved yields, but also showed less thermal degradation in the final extracts. Several studies have been conducted into the extraction of the main aroma compounds from spices. For example, vanillin was extracted from vanilla pods (Jadhav et al., 2009) carvone from caraway seeds (Chemat et al., 2004) and safranal from Greek saffron (Kanakis et al., 2004). Yields of vanillin obtained after 1 h using UAE were similar to those obtained after 8 h using conventional extraction methods. Wines and other alcoholic beverages contain many volatile compounds, which play an important role in organoleptic characteristics. The ultrasound-assisted maceration of fruits and herbs in alcoholic beverages to extract volatile compounds is an important process in the food industry (Fig. 3.12). Antioxidants A wide variety of fruits and vegetables have been studied by UAE because antioxidants are present in different amounts in different varieties of plants, and these antioxidants come from different families. One of the most common antioxidants is the lycopene extracted from tomatoes (Lianfu and Zelong, 2008). In this particular work, authors not only worked on UAE, but also on coupling ultrasounds with microwaves, which produces high lycopene extraction in only 6 min. Herrera and Luque de Castro (2005) showed that the amount of phenolic compounds that could be extracted from strawberries in 2 min, was similar to the amount extracted over 20 h using the conventional method, and 3 h using supercritical fluid extraction. Another example is the extraction of anthocyanins from raspberries, developed by Chen et al. (2007). In this work, the same yields of anthocyanins were achieved in 3 min when using UAE, compared with 53 min when using the conventional extraction system. The difference between ultrasound and conventional methods will be more or less significant, depending on the part of the plant studied. This has been highlighted by studies of grapes and grape seeds (Palma and Barroso, Sono-extraction
0.100
Conventional extraction
Yield (%)
0.075 0.050
0.025 0.000
0
30
60
90
120 150 180 210 240 270 Time (min)
Fig. 3.12 Rapid sono-extraction of orange peels in alcoholic beverages.
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Principles of physically assisted extractions and applications 95 2002). In this work, authors found that the use of ultrasound instead of maceration was more interesting when working on seeds than on the whole fruit (Table 3.1). Oil and fat Many papers have reported on the UAE of oil and fat from various food samples. According to Luque-Garcia and Luque de Castro (2004), extraction from oleaginous seeds is difficult. Indeed, only 75–85% of the oil is solubilised in the solvent. The rest of the oil content is strongly bound to the matrix and cannot be extracted without additional treatments. For instance, Li et al. (2004) described the UAE of soybean oil in an optimised yield and reduced operating time, compared with conventional maceration. Luque-Garcia and Luque de Castro (2004) have reported a device consisting of an ultrasoundassisted Soxhlet apparatus, which is an attractive alternative to traditional Soxhlet extraction. Using this device, it takes only 90 minutes to obtain the same yield (99% of fat recovery) as can be obtained over 12 h with the traditional Soxhlet extraction method (Table 3.1). 3.5.5 Hazard analysis critical control point (HACCP) for UAE processing operation The hazard analysis critical control point (HACCP) system is a process that identifies and assesses the hazards and risks associated with the manufacture, Table 3.1 Ultrasound-assisted extraction of food ingredients Matrix (reference)
Extracts
Essential oils and aromas Citrus flowers (Alissandrakis et al., 2003) Caraway seeds (Chemat et al., 2004) Garlic (Kimbaris et al., 2006) Lavender flowers (Da Porto et al., 2009) Greek saffron (Kanakis et al., 2004) Antioxidant extracts Strawberries (Herrera and Luque de Castro, 2005) Raspberries (Chen et al., 2007) Tomatoes (Liangfu and Zelong, 2008) Grape (Palma and Barroso, 2002) Fat and oil extraction Olive seeds (Luque-Garcia and Luque de Castro, 2004) Soybean germ and seaweed (Li et al., 2004) Cocoa powder and nibs (Luque- Garcia and Luque de Castro, 2004) Bakery products (Luque-Garcia and Luque de Castro, 2004)
Linalool Carvone and limonene Essential oil (organosulfur compounds) Essential oil (1,8-cineol, camphor and linalyl acetate) Safranal Polyphenolic compounds Anthocyanins Lycopene Phenolic Oil Oil Fat content Fat content
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96 Separation, extraction and concentration processes storage and distribution of foods and implement the appropriate controls, aiming at the elimination or reduction of these hazards at specific points of the production line. In UAE processing, parameters such as the temperature and physical properties of the product, the timescale of treatment in the ultrasound chamber, the frequency and power of the ultrasound, and the nature of the probe are points which need to be monitored. For each part of the food process, including ultrasound processing, the evaluation and classification of hazards are completed. In this way, a grade is given to each hazard that has been identified according to its gravity, its risk of occurrence and the ability to detect it. For each hazard, preventive measures are set up, with procedures that indicate who will be responsible for them, as well as how, when and where. Preventive measures are aimed primarily at avoiding the occurrence of the hazard. For example, a good maintenance plan decreases the risk of metal contamination from the ultrasound horn (Table 3.2).
3.6 Microwave-assisted extraction (MAE) in the food industry 3.6.1 Principles of MAE Microwaves are electromagnetic waves with a frequency range of between 100 MHz and 3 GHz that comprise electric and magnetic field components and thus constitute propagating electromagnetic energy. This energy acts as a non-ionising radiation, which causes molecular motion of ions and rotation of the dipoles, but does not affect molecular structure. Table 3.2 Possible critical control point limits and associated corrective actions in an ultrasound-assisted extraction processing Critical control point
Danger
Target
Deviation
Processing Temperature Adequate Does not conform operation to the requirements Ultrasound Adequate Does not conform frequency to the requirements Ultrasound Adequate Does not conform power to the requirements Flow rate Below Does not conform specifications to the requirements Probe Absence Detected spoilage
Corrective actions
Adjust temperature reject or reprocess the product Adjust frequency, laboratory controls, or reprocess the product Adjust the power, laboratory controls, or reprocess the product Adjust flow, laboratory controls, or reprocess the product Reject
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Principles of physically assisted extractions and applications 97 When dielectric materials containing either permanent or induced dipoles are placed in a MW field, the rotation of the dipoles in the alternating field produces heat. More precisely, the applied MW field causes the molecules to spend slightly more time, on average, orienting themselves in the direction of the electric field rather than in other directions. When the electric field is removed, thermal agitation returns the molecules to a disordered state in the relaxation time and thermal energy is released. Thus, MW heating results from the dissipation of the electromagnetic waves in the irradiated medium (Fig. 3.13). The amount of dissipated power in the medium depends on the complex permittivity of the material and the local time-averaged electric field strength (Metaxas and Meredith, 1993). In conventional heating, heat is transferred from the heating medium to the interior of the sample, whereas in MW heating; heat is dissipated volumetrically inside the irradiated medium, and thus heat transfers occur from the sample to the colder environment. This causes a significant difference Moisture (%)
Classical heating
Microwave heating
Microwave heating Heat transfer
1 2
1. Solid friction desorption 2. Internal diffusion 3. External diffusion
Mass transfer
3
Conventional heating Heat transfer Mass transfer
Fig. 3.13 Microwave versus conventional extraction.
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98 Separation, extraction and concentration processes between conventional and MW heating. In conventional heating, heat transfer depends on thermal conductivity, on the temperature difference across the sample, and, for fluids, on convection currents. As a result, the temperature increase is often rather slow. In contrast, the volumetric heating effect in MW heating enables much faster temperature increases to be obtained, depending on the MW power and the dielectric loss factor of the material being irradiated. The influence of MW energy on chemical or biochemical reactions is strictly thermal. The MW energy quantum is given by the standard equation W = hn. Within the frequency domain of microwaves and hyper-frequencies (300 MHz–300 GHz), the corresponding energies are respectively 1.24 ¥ 10–6 eV to 1.24 ¥ 10–3 eV. These energies are much lower than the usual ionisation energies of biological compounds (13.6 eV), of covalent bond energies like OH (5 eV), hydrogen bonds (2 eV), and van der Waals intermolecular interactions (lower than 2 eV) and even lower than the energy associated with Brownian motion at 37 °C (2.7 ¥ 10–3 eV). From a scientific point of view, the direct molecular activation of microwaves should be excluded. 3.6.2 MAE reactors It is now possible to use a wide range of vessels and instrumentation when working with microwaves, depending on the intended purpose and the analytes to be extracted. Two types of MW ovens can be found in laboratories and they differ in terms of their cavity type: single-mode (also called monomode or focused mode) and multi-mode. A single-mode cavity permits the generation of one type of frequency, allowing only one mode of resonance to be excited. Thus, the sample is placed in such a way as to receive the maximum amount of MW energy, ensuring a direct and efficient application of a high density of MW energy to the sample matrix, and inducing a rapid heating of the medium with a fast extraction of the analyte. However, monomode systems are quite limited in terms of the volume and quantity of the sample that is ultimately useable. On the other hand, multi-mode ovens are larger and allow several modes of resonance. Thus, a large number of incident waves can be applied to the sample being studied. This kind of oven allows for a large quantity of matrix, and can be used with several types and forms of vessels. There are also MW reactors conceived especially for solvent extraction in the laboratory (0.1 to 1 l) and on an industrial scale (100 l h–1) (Fig. 3.14). 3.6.3 MAE: main applications Microwaves are very useful for rapid extraction and obtaining high quality aromatic compounds from garlic, lavender flowers, orange peel and rosemary leaves. The extraction of essential oils from orange peels with solvent-free MW extraction is better in terms of energy saving, extraction time (30 min
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Principles of physically assisted extractions and applications 99
(a)
(b) –1
Fig. 3.14 (a) Laboratory (1 l) and (b) industrial (100 l h ) MAE reactors (with permission, www.milestonesrl.com and www.rchimex.com).
versus 3 h), product yield (0.42 versus 0.39%) and product quality (AbertVian et al., 2008). Hemwimon et al. (2007) made a comparison between the MAE of antioxidative anthraquinones from the roots of Morinda citrifolia, and extraction with other conventional techniques (maceration and Soxhlet). The efficiency of extraction using MAE (15 min) was much higher than that using maceration (3 days) and was also similar to that of Soxhlet (4 h). Similarly, the best recoveries of antioxidants in short time periods, and with a lower consumption of polar solvent were also observed from sweet grass by Grigonis et al. (2005). As well as obtaining a higher yield, Pan et al. (2008) extracted phenolic compounds with scavenging ability, compared with synthetic antioxidants from longan peel with MAE. Elkhori et al. (2007) described the MAE of cocoa powder with hexane/ isopropanol, which resulted in the rapid determination of fat contents, with recoveries similar to, or better than, those of the conventional method, and with the added advantages of low solvent consumption, short extraction time, low energy consumption and excellent reproducibility. Focused MW-assisted Soxhlet extraction was used by Pérez-Serradilla et al. (2007) for acorn oil determination, which resulted in trans fatty acid-free oil, probably because of the reduced exposure to drastic conditions in just 30 min, which is much less than the time required by the Soxhlet (8 h) and stirring (56 h) reference methods. Lianfu and Zelong (2008) described the extraction of lycopene from tomatoes with the combined innovative techniques of ultrasound and MW-assisted extraction (UMAE) in comparison with ultrasound-assisted
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100 Separation, extraction and concentration processes extraction (UAE). Similarly, a higher yield of anthocyanins in red raspberries was also observed by Sun et al. (2007) with optimal conditions of MAE. MAE proved to be more rapid and efficient as it extracts various types of anthocyanins, without any destruction of their chemical structure, in a very short time, owing to the intensive disruption of tissue structure under MW irradiation. Table 3.3 clearly shows some applications of MAE for various useful metabolites, carried out using several vegetal matrices. 3.6.4 Safety considerations The MAE process is simple and can be readily understood in terms of the operating steps to be performed. However, in inexperienced hands, the application of MW energy can pose serious hazards. All persons dealing with microwaves must exercise a high level of safety and attention to detail when planning and performing such experiments. They have to ensure that they seek proper information from knowledgeable sources and that they do not attempt to implement this technique unless proper guidance is provided. Only approved equipment and scientifically sound procedures should be used.
3.7 Combination of physical treatments for extraction in the food industry Various physical treatments can induce various effects on cell and tissue structure, which may be more or less beneficial for the extraction of intracellular components. Unfortunately, the studies comparing or combining two or more Table 3.3 Microwave-assisted extraction of food ingredients Matrix (reference)
Extracts
Antioxidant extracts Green tea leaves (Pan et al., 2003) Noni plant roots (Hemwimon et al., 2007) Sea buckthorn (Sharma et al., 2008) Longan peel (Pan et al., 2008)
Polyphenols Anthraquinones Phenolic constituents Flavonoids
Fat and oil extraction Olive seeds (Virot et al., 2007) Oleaginous seeds (Pérez-Serradilla et al., 2007) Soybean germ and seaweed (Cravotto et al., 2008) Cocoa powder and nibs (Elkhori et al., 2007)
Oil Oil Oil Fat content
Natural food colour extraction Tomato paste (Lianfu and Zelong, 2008) Curcuma (Mandal et al., 2008) Red raspberries (Sun et al., 2007) Paprika (Csiktusnádi Kiss et al., 2000)
Lycopene Curcumin Anthocyanins Carotenoids
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Principles of physically assisted extractions and applications 101 physically assisted extractions are rather rare. El-Belghiti (2005) compared the application of PEF (680 V cm–1, 1000 pulses of 100 ms) and HVED (40 kV, 100 discharges) for the cold (20 °C) aqueous extraction of solutes from two dried products: tea leaves and Datura innoxia (common name: moonflower) roots. PEF treatment did not have any influence on the process of extracting from dried and rehydrated products, probably because of the absence of the cell membrane element, which had been disrupted by previous drying. The non-efficacy of moderate electric treatment for the dried cell tissue was reported earlier in Sensoy and Sastry (2004). In contrast to PEF, the HVED smashed products to small fragments and the extraction kinetics were significantly enhanced for both tea leaves and Datura innoxia. El-Belghiti et al. (2007) compared PEF, HVED and PU as treatments to enhance cold (20 °C) aqueous extraction from fennel gratings. The objective was to obtain extracts used as natural food preservatives (antioxidants). The three assisted extractions led to the same final yield of solutes, and all preserved the antioxidant substances. However, their kinetics were different. The most rapid and the slowest kinetics were offered by HVED and PU, respectively. The final yield of 98% was reached in 20 min with HVED, in 40 min with PEF, and in more than 180 min with PU. The amounts of energy needed for these treatments were also different: the PEF treatment appeared to be the most economic energetically. Recently, Shynkaryk et al. (2008) compared PEF, HVED, and high-pressure homogenisation (HPH) for disruption of wine yeast cells (S. cerevisiae bayanus). The PEF and HVED were applied at electric field strengths of 10 kV cm–1 and 40 kV cm–1, respectively. The HPH was applied within a pressure range of 30–200 MPa for a different number of passes (1–20). It was shown that the releasing efficiency of ionic components, enzymes, proteins and other bioproducts dramatically depended on the applied method of disruption. Extraction efficiency was rather high for ionic components and small for high molecular weight components in PEF or HVED pretreated suspensions. The PEF treatment removes membrane barriers and accelerates the release of ionic contents, but has practically no influence on the cell walls. The efficiency of releasing high molecular weight content is questionable and may depend on the cell strain, the age of culture, the time of extraction, temperature and many other factors. For example, high levels of PEF extraction of proteins for baker’s yeast was previously reported (Ganeva et al., 2003), but the efficiency of the PEF extraction of proteins was found to be rather low for wine yeast in the study by Shynkaryk et al. (2008). On the contrary, the application of the combined HVED–HPH technique resulted in a material loss of integrity in the yeast cell walls. Even a small disintegration (Zi = 0.15±0.05), initially induced by HVED pretreatment, resulted in noticeable acceleration of HPH disruption kinetics. Although most of the research effort in UAE has concentrated on ultrasound itself, some studies have also examined the coupling between ultrasound and other techniques. For instance, UAE is currently being employed in
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102 Separation, extraction and concentration processes combination with MW energy (Lagha et al., 1999), supercritical fluid extraction (Hu, 2007) or simply with conventional methods such as Soxhlet extraction (Luque-García and Luque de Castro, 2004). When combined with supercritical fluid extraction, UAE enhances the mass transfer of the species of interest, from the solid phase to the solvent used for extraction. Soxhlet extraction can also be improved by ultrasound when applied at the cartridge zone before siphoning, thus permitting the removal of lipid fractions from very compact matrices. The efficiency of combining MW and ultrasound has been clearly shown in applications such as the extraction of food ingredients. Studies combining assisted extractions present undeniable interest as synergistic effects of different treatments may be discovered.
3.8 References Abert-Vian M, Fernandez X, Chemat F (2008), ‘Microwave hydrodiffusion and gravity, a new technique for extraction of essential oils’, Journal of Chromatography A, 1190, 14–18. Alissandrakis E, Daferera D, Tarantilis P A, Polissiou M and Harizanis P C (2003), ‘Ultrasound-assisted extraction of volatile compounds from citrus flowers and citrus honey’, Food Chemistry, 82, 575–579. Angersbach A, Heinz V and Knorr D (2002), ‘Evaluation of process-induced dimensional changes in the membrane structure of biological cells using impedance measurement’, Biotechnology Progress, 18(3), 597–603. Asfaw N, Licence P, Novitskii AA, Poliakoff M (2005), ‘Green chemistry in Ethiopia: the cleaner extraction of essential oils from Artemisia afra: a comparison of clean technology with conventional methodology’, Green Chemistry, 7, 352–356. Barbosa-Cánovas G V, Pothakamury U R, Palou E and Swanson B G (1998), Nonthermal preservation of foods, Marcel Dekker, New York. Barskaya A V, Kuretz B I and Lobanova G L (2000), ‘Extraction of water soluble matters from vegetative raw material by electrical pulsed discharges’, 1st International Congress on Radiation Physics, High Current Electronics, and Modification of Materials, Tomsk, Russia, 533–535. Bazhal M and Vorobiev E (2000), ‘Electrical treatment of apple cossettes for intensifying juice pressing’, Journal of the Science of Food and Agriculture, 80, 1668–1674. Bazhal M I, Lebovka NI and Vorobiev E I (2001), ‘Pulsed electric field treatment of apple tissue during compression for juice extraction’, Journal of Food Engineering 50, 129–139. Biss C H, Combes S A and Skudder P J (1989), ‘The development and application of ohmic heating for the continuous processing of particulate food stuffs’, In Processing engineering in the food industry (Field RW, Howell JA, eds), Elsevier, London, 17–27. Bluhm H (2006), Pulsed power systems, Springer, Berlin, Heidelberg, New York. Bluhm H and Sack M (2008), ‘Industrial-scale treatment of biological tissues with pulsed electric field’, In: E. Vorobiev and N. Lebovka (Eds), Electrotechnologies for extraction from food plants and biomaterials, Springer, 237–269. Boussetta N, Lebovka N, Vorobiev E, Adenier H, Bedel-Cloutour C, and Lanoisellé J-L (2009a), ‘Electrically assisted extraction of soluble matter from Chardonnay grape skins for polyphenols recovery’, Journal of Agricultural and Food Chemistry, 57, 1491–1497.
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Principles of physically assisted extractions and applications 103 Boussetta N, Lanoisellé J-L, Bedel-Cloutour C and Vorobiev E (2009b), ‘Extraction of polyphenols from grape pomace by high voltage electrical discharges: effect of sulphur dioxide, freezing process and temperature’ Journal of Food Engineering, 95, 192–198. Bouzrara H (2001), Amélioration du pressage de produits végétaux par Champ Electrique Pulsé. Cas de la betterave à sucre, PhD thesis, UTC, Compiègne, France. Bouzrara H and Vorobiev E (2000), ‘Beet juice extraction by pressing and pulsed electric fields’, International Sugar Journal, CII(1216), 194–200. Bouzrara H and Vorobiev E (2003), ‘Solid/liquid expression of cellular materials enhanced by pulsed electric field’, Chemical Engineering and Processing, 42, 249–257. Butz P and Tauscher B (2002), Emerging technologies: chemical aspects, Food Research International, 35, 279–284. Chemat S, Lagha A, AitAmar H, Bartels P, Chemat F (2004), ‘Comparison of conventional and ultrasound-assisted extraction of carvone and limonene from caraway seeds’, Flavour and fragrance journal, 19, 188–193. Chen F, Sun Y, Zhao G, Liao X, Hu X, Wu J, Wang Z (2007), ‘Optimization of ultrasoundassisted extraction of anthocyanins in red raspberries and identification of anthocyanins in extract using high-performance liquid chromatography-mass spectrometry’, Ultrasonics Sonochemistry, 14, 767–772. Cravotto G, Boffa L, Mantegna S, Perego P, Avogadro M, Cintas P (2008), ‘Improved extraction of vegetable oils under high-intensity ultrasound and/or microwaves’, Ultrasonics Sonochemistry, 15, 898–902. Csiktusnádi Kiss G, Forgács E, Cserháti T, Mota T, Morais H, Ramos A (2000), ‘Optimisation of microwave-assisted extraction of pigments from paprika powders’, Journal of Chromatography A, 889, 41–49. Da Porto C, Decorti D and Kikic I (2009), ‘Flavour compounds of Lavandula angustifolia L., to use in food manufacturing: comparison of three different extraction methods’, Food Chemistry, 112, 1072–1076. De Gennaro L, Cavella S, Romano R and Masi P (1999), ‘The use of ultrasound in food technology I: inactivation of peroxidase by thermosonication’, Journal of Food Engineering 39, 401–407. Dugand L R (1980), ‘Natural antioxidants’, In M.G. Simic, M. Karel (Eds), Autooxidation in food and biological systems, Plenum Press, New York, 261–295. El-Belghiti K and Vorobiev E (2004), ‘Mass transfer of sugar from beets enhanced by pulsed electric field’, Transactions of the Institution of Chemical Engineers, 82, 226–230. El-Belghiti K (2005), Effets d’un champ électrique pulsé sur le transfert de matière et sur les caractéristiques végétales, PhD thesis, UTC, Compiègne, France. El-Belghiti K, Moubarik R and Vorobiev E (2007), Use of moderate pulsed electric field, electrical discharges and ultrasonic irradiations to improve aqueous extraction of solutes from fennel (Foeniculum vulgare), unpublished data. El-Belghiti K and Vorobiev E (2005a), ‘Kinetic model of sugar diffusion from sugar beet tissue treated by pulsed electric field’, Journal of the Science of Food and Agriculture, 85, 213–218. El-Belghiti K and Vorobiev E (2005b), ‘Modelling of solute aqueous extraction from carrots subjected to a pulsed electric field pre-treatment’, Biosystems Engineering, 90(3), 289–294. Elkhori S, Jocelyn Paré J, Bélanger J, Pérez E (2007), ‘The microwave-assisted process: extraction and determination of fat from cocoa powder and cocoa nibs’, Journal of Food Engineering, 79, 1110–1114. El Zakhem H, Lanoisellé J-L, Lebovka N I, Nonus M and Vorobiev E (2006a), ‘Behavior of yeast cells in aqueous suspension affected by pulsed electric field’, Journal of Colloid and Interface Science, 300(2), 553–563. El Zakhem H, Lanoisellé J-L, Lebovka N I, Nonus M and Vorobiev E (2006b), ‘The early
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104 Separation, extraction and concentration processes stages of Saccharomyces cerevisiae yeast suspensions damage in moderate pulsed electric fields’, Colloids and Surfaces B: Biointerfaces, 479(2), 189–197. Eshtiaghi, MN and Knorr D, (2002), ‘High electric field pulse pretreatment: Potential for sugar beet processing’, Journal of Food Engineering, 52, 265–272. Fincan M, DeVito F and Dejmek P (2004), ‘Pulsed electric field treatment for solid–liquid extraction of red beetroot pigment’, Journal of Food Engineering, 64(3), 381–388. Fincan M and Dejmek P (2002), ‘In situ visualization of the effect of a pulsed electric field on plant tissue’, Journal of Food Engineering, 55, 223–230. Ganeva V, Galutzov B and Teissié J (2003), ‘High yield electroextraction of proteins from yeast by a flow process’, Analytical Biochemistry 315(1), 77–84. Grémy-Gros C, Lanoisellé J-L and Vorobiev E (2008) ‘Application of high-voltage electrical discharges for the aqueous extraction from oilseeds and other plants’, In: E. Vorobiev and N. Lebovka (Eds), Electrotechnologies for extraction from food plants and biomaterials, Springer, New York, 217–236. Grigonis D, Venskutonis P, Sivik B, Sandahl M, Eskilsson C (2005), ‘Comparison of different extraction techniques for isolation of antioxidants from sweet grass’, Journal of Supercritical Fluids, 33, 223–233. Grimi N, Lebovka M, Vorobiev E, Vaxelaire J (2009a), ‘Compressing behaviour and texture evaluation for potatoes pretreated by pulsed electric field’, Journal of Texture Studies, 40, 208–224. Grimi N, Lebovka M, Vorobiev E and Vaxelaire J (2009b), ‘Effect of a pulsed electric field treatment on expression behaviour and juice quality of Chardonnay grape’, Food Biophysics, 4, 1557–1858. Grimi N, Praporscic I, Lebovka N and Vorobiev E (2007), ‘Selective extraction from carrot slices by pressing and washing enhanced by pulsed electric fields’, Separation and Purification Technology, 58(2), 267–273. Grimi N, Vorobiev E and Vaxelaire J (2008), ‘Juice extraction from sugar beet slices by belt filter press: effect of pulsed electric field and operating parameters’, 14th World Congress of Food Science and Technology, 19–23 October 2008, Shanghai, China, TS24–29, 554. Gros C, Lanoisellé J-L and Vorobiev E (2003), ‘Towards an alternative extraction process for linseed oil’, Chemical Engineering Research and Design, 81(9), 1059–1065. Hemwimon S, Pavasant P and Shotipruk A (2007), Microwave-assisted extraction of antioxidative anthraquinones from roots of Morinda citrifolia, Separation and Purification Technology, 54, 44–50. Herrera M and Luque de Castro M D (2005), Ultrasound-assisted extraction of phenolic compounds from strawberries prior to liquid chromatographic separation and photodiode array ultraviolet detection, Journal of Chromatography A, 1100, 1–8. Hu A (2007), ‘Ultrasound assisted supercritical fluid extraction of oil and coixenolide from adlay seed’, Ultrasonics Sonochemistry, 14, 219–222. Imai T, Uemura K, Ishida N, Yoshizaki S, and Noguchi A (1995), ‘Ohmic heating of Japanese white radish Rhaphanus sativus L’, International Journal of Food Science and Technology, 30(4), 461–472. Jadhav D, Rekha B N, Gogate P R, Rathod V K (2009), ‘Extraction of vanillin from vanilla pods: A comparison study of conventional soxhlet and ultrasound-assisted extraction’, Journal of Food Engineering, 93, 421–426. Jaeger H, Balasa A, and Knorr D (2008), ‘Food industry applications for pulsed electric fields’, In: E Vorobiev and N Lebovka (Eds), Electrotechnologies for extraction from food plants and biomaterials, Springer, 181–216. Jemai A B and Vorobiev E (2002), ‘Effect of moderate electric field pulse (MEFP) on the diffusion coefficient of soluble substances from apple slices’, International Journal of Food Science and Technology, 37, 73–86. Jemai A B and Vorobiev E (2003), ‘Enhancing leaching from sugar beet cossettes by pulsed electric field’, Journal of Food Engineering, 59, 405–412.
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Principles of physically assisted extractions and applications 105 Jemai A B and Vorobiev E (2006), ‘Pulsed electric field assisted pressing of sugar beet slices: towards a novel process of cold juice extraction’, Biosystems Engineering, 93(1), 57–68. Kanakis C D, Daferera D J, Tarantilis P, Polissiou, M (2004), ‘Qualitative determintaion of volatile compounds and quantitative evaluation of safranal and 4-hydroxy-2,6,6,trimethyl-1-cyclohexane-1-carboxaldehyde (HTCC) in Greek saffron’, Journal of Agriculture and Food Chemistry, 52, 4515–4520. Kemp M R, and Fryer P J (2007), ‘Enhancement of diffusion through foods using alternating electric fields’, Innovative Food Science and Emerging Technologies, 8, 143–153. Kimbaris AC, Siatis N G, Daferera D J, Tarantilis P A, Pappas C S and Polissiou M G (2006), ‘Comparison of distillation and ultrasound-assisted extraction methods for the isolation of sensitive aroma compounds from garlic’, Ultrasonics Sonochemistry, 13, 54–59. Knorr D, Angersbach A, Eshtiaghi M N, Heinz V and Lee D-U (2001), ‘Processing concepts based on high intensity electric field pulses’, Trends in Food Science and Technology 12(3–4), 129–135. Knorr D, Geulen M, Grahl, T and Sitzmann W (1994), ‘Food application of high electric field pulses’, Trends in Food Science and Technology, 5, 71–75. Konduser M and Miklavcic D (2008), ‘Electroporation in biological cell and tissue: an overview’ In: E Vorobiev and N Lebovka (Eds), Electrotechnologies for extraction from food plants and biomaterials, Springer, New York, 1–38. Kulshrestha S and Sastry S K (2003), ‘Frequency and voltage effects on enhanced diffusion during moderate electric field (MEF) treatment’, Innovative Food Science and Emerging Technologies, 4, 189–194. Lagha A, Chemat S, Bartels P and Chemat F (1999), ‘Microwave–ultrasound combined reactor suitable for atmospheric sample preparation procedure of biological and chemical products’, Analusis, 27, 452–455. Lazarenko B R, Fursov S P, Scheglov Y A, Bordiyan V V, Chebanu V G (1977), Electroplasmolysis, Karta Moldavaneske, Kishinev, USSR (in Russian). Lebovka N I, Bazhal M I and Vorobiev E (2000), ‘Simulation and experimental investigation of food material breakage using pulsed electric field treatment’, Journal of Food Engineering 44, 213–223. Lebovka N I, Bazhal M I and Vorobiev E (2001), ‘Pulsed electric field breakage of cellular tissues: visualization of percolative properties’, Innovative Food Science and Emerging Technologies, 2, 113–125. Lebovka N I, Praporscic I and Vorobiev E (2004a), ‘Combined treatment of apples by pulsed electric fields and by heating at moderate temperature’, Journal of Food Engineering, 65, 211–217. Lebovka N I, Praporscic I and Vorobiev E (2004b), ‘Effect of moderate thermal and pulsed electric field treatments on textural properties of carrots, potatoes and apples’, Innovative Food Science and Emerging Technologies, 5, 9–16. Lebovka N I, Praporscic I, Ghnimi S and Vorobiev E (2005), ‘Does electroporation occur during the ohmic heating of food?’, Journal of Food Science, 70(5), E308–311. Lebovka N I, Shynkaryk M and Vorobiev E (2007), ‘Moderate electric fields treatment of sugarbeet tissues’, Biosystems Engineering, 96(1), 47–56. Li H, Pordesimo L and Weiss J (2004), High intensity ultrasound-assisted extraction of oil from soybeans, Food Research International, 37, 731–735. Li L, Ding L, Vorobiev E and Lanoiselle J-L (2009), ‘The application of high voltage electrical discharge for oil extraction from oil meals of linseed, rapeseed and palm kernel in aqueous solution’, Conference on Food Engineering CoFE 2009, April 5–8, 2009, Columbus, USA. Lianfu Z and Zelong L (2008), Optimization and comparison of ultrasound/microwave assisted extraction (UMAE) and ultrasonic assisted extraction (UAE) of lycopene from tomatoes, Ultrasonics Sonochemistry, 15, 2008, 731–737.
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106 Separation, extraction and concentration processes Lima M, Heskitt B and Sastry S (1999), ‘The effect of frequency and wave form on the electrical conductivity–temperature profiles of turnip tissue’, Journal of Food Process Engineering, 22(1), 41–54. Lima M, Heskitt B F and Sastry S K (2001), ‘Diffusion of beet dye during electrical and conventional heating at steady-state temperature’, Journal of Food Process Engineering, 24, 331–340. Lopez N, Puertolas E, Condon S, Raso J and Alvarez I (2009a), ‘Enhancement of the extraction of betanine from red beetroot by pulsed electric fields’, Journal of Food Engineering, 90(1), 60–66. Lopez N, Puertolas E, Hernández-Orte P, Alvarez I and Raso J (2009b), ‘Effect of a pulsed electric field treatment on the anthocyanins composition and other quality parameters of Cabernet Sauvignon freshly fermented model wines obtained after different maceration times’, LWT – Food Science and Technology, 42(7), 1225–1231. Luque-García J L and Luque de Castro M D (2004), ‘Ultrasound-assisted Soxhlet extraction: an expeditive approach for solid sample treatment: application to the extraction of total fat from oleaginous seeds’, Journal of Chromatography A, 1034, 237–241. Mandal V, Mohan Y and Hemalatha S (2008), ‘Microwave assisted extraction of curcumin by sample-solvent dual heating mechanism using Taguchi L9 orthogonal design’, Journal of Pharmaceutical and Biomedical Analysis, 46, 322–327. Mason T J (1990), Chemistry with ultrasound, Elsevier Applied Science, New York. McLellan M R, Kime R L and Lind K R (1991), ‘Electroplasmolysis and other treatments to improve apple juice yield’, Journal of Science of Food and Agriculture, 57(2), 303–306. Metaxas A C and Meredith R J (1993), Industrial microwave heating, IEEE, London. Palma M and Barroso C (2002), ‘Ultrasounds-assisted extraction and determination of tartaric and malic acids from grapes and winemaking by-products’, Analitica chimica acta, 458, 119–125. Pan X, Niu G, Liu H (2003), ‘Microwave-assisted extraction of tea polyphenols and tea caffeine from green tea leaves’, Chemical Engineering and Processing, 42, 129–133. Pan Y, Wang K, Huang S, Wang H, Mu X, He C, Ji X, Zhang J, Huang F (2008), ‘Antioxidant activity of microwave-assisted extract of longan peel’, Food Chemistry, 106, 1264–1270. Pérez-Serradilla J, Ortiz M, Sarabia L, Luque de Castro M (2007), ‘Focused microwaveassisted soxhlet extraction of acorn oil for determination of the fatty acid profile by GC-MS’, Analytical and Bioanalytical Chemistry, 388, 451–462. Povey M and Mason T J (1998), Ultrasound in Food Processing, Blackie Academic & Professional, London. Praporscic I, Ghnimi S and Vorobiev E (2005), ‘Enhancement of pressing of sugar beet cuts by combined pulsed electric field and ohmic heating’, Journal of Food Processing and Preservation, 29(5-6), 378–389. Praporscic I, Lebovka N I, Ghnimi S and Vorobiev E (2006), ‘Ohmically heated, enhanced expression of juice from apple and potato tissues’, Biosystems Engineering, 93(2), 199–204. Praporscic I, Lebovka NI, Vorobiev E, Mietlon-Peuchot M (2007), ‘Pulsed electric field enhanced expression and juice quality of white grapes’, Separation and Purification Technology, 52(3), 520–526. Praporscic I, Shynkaryk M, Lebovka N and Vorobiev E (2007b), ‘Analysis of juice colour and dry matter content during pulsed electric field enhanced expression of soft plant tissues’, Journal of Food Engineering, 79(2), 662–670. Priego-Capote F and Luque de Castro M D (2007), Ultrasound-assisted digestion: a useful alternative in sample preparation, Journal of Biochemical and Biophysical Methods, 70, 299–310.
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Principles of physically assisted extractions and applications 107 Raso J, Manas P, Pagan R and Sala FJ (1999), ‘Influence of different factors on the output power transferred into medium by ultrasound’, Ultrasonics Sonochemistry, 5, 157–162. Rogov I A and Gorbatov A V (1974), Physical methods of foods processing, Pischevaja Promyshlennost, Moscow (in Russian). Sastry S K (2005), ‘Advances in ohmic heating and moderate electric field (MEF) Processing’, In: G Barbosa-Canovas, M S Tapia and M P Cano (Eds), Novel food processing technologies, CRC Press, New York, 491–499. Schilling S, Alber T, Toepfl S, Neidhart S, Knorr D, Schieber A and Carle R (2007), ‘Effects of pulsed electric field treatment of apple mash on juice yield and quality attributes of apple juices’, Innovative Food Science and Emerging Technologies, 8, 127–134. Schreier P J R, Reid D G and Fryer P J (1993), ‘Enhanced diffusion during the electrical heating of foods’, International Journal of Food Science and Technology, 28, 249–260. Schrive L, Grasmick A, Moussière S, Sarrade S, (2006), ‘Pulsed electric field treatment of Saccharomyces cerevisiae suspensions: a mechanistic approach coupling energy transfer, mass transfer and hydrodynamics’, Biochemical Engineering Journal, 27, 212–224. Sensoy I and Sastry S K (2004), ‘Extraction using moderate electric fields’, Journal of Food Science: Food Engineering and Physical Properties, 69(1), 7–13. Sharma U, Singh H, Sinha A (2008), ‘Microwave-assisted efficient extraction of different parts of Hippophae rhamnoides for the comparitive evaluation of antioxidant activity and quantification of its phenolic constituents by reverse-phase high-performance liquid chromatography (RP-HPLC)’, Journal of Agriculture Food Chemistry, 56, 374–379. Shotipruk A, Kaufman P B and Wang H Y (2001), ‘Feasability study of repeated harvesting of menthol from biologically viable Mentha x piperata using ultrasonic extraction’, Biotechnology progress, 17, 924–928. Shynkaryk M V, Lebovka N I and Vorobiev E (2008), ‘Pulsed electric fields and temperature effects on drying and rehydration of red beetroots’, Drying Technology, 26(6), 695–704. Shynkaryk M V, Lebovka N I, Lanoisellé J-L, Nonus M, Bedel-Clotour C and Vorobiev E (2009), ‘Electrically-assisted extraction of bio-products using high pressure disruption of yeast cells (Saccharomyces cerevisiae)’, Journal of Food Engineering, 92( 2), 189–195. Sugiarto A T, Sato M and Skalny J D (2001) ‘Transient regime of pulsed breakdown in low-conductive water solutions’, Journal of Physics D: Applied Physics, 34(23), 3400–3406. Sun Y, Liao X, Wang Z, Hu X, Chen F (2007), ‘Optimization of microwave-assisted extraction of anthocyanins in red raspberries’, European Food Research Technology, 225, 511–523. Suslick K S (1988), ‘Ultrasound, its chemical, physical and biological effects’, VCH Publishers, New York. Van der Poel P W, Schiweck H and Schwartz T (1998), Sugar technology beet and cane sugar manufacture, Beet Sugar Development Foundation. Denver, USA. Virot M, Tomao V, Colnagui G, Visinoni F, Chemat F (2007), ‘New microwave-integrated Soxhlet extraction an advantageous tool for the extraction of lipids from food products’, Journal of Chromatography A, 1174, 138–144. Vorobiev E, Jemai A B, Bouzrara H, Lebovka N I and Bazhal M I (2005), ‘Pulsed electric field assisted extraction of juice from food plants’, In: G Barbosa-Canovas, M S. Tapia and M P Cano (Eds), Novel food processing technologies, CRC Press, New York, 105–130. Vorobiev E and Lebovka N I (2006), ‘Extraction of intercellular components by pulsed
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108 Separation, extraction and concentration processes electric fields’, In: J Raso and V Heinz (Eds), Pulsed electric field technology for the food industry. Fundamentals and applications, Springer, New York, 153–194. Vorobiev E and Lebovka N I (2008), ‘Pulsed electric field induced effects in plant tissues: fundamental aspects and perspectives of application’. In: E Vorobiev and N Lebovka (Eds), Electrotechnologies for extraction from food plants and biomaterials, Springer, New York, 39–82. Wang W C and Sastry S K (2002), ‘Effects of moderate electrothermal treatments on juice yield from cellular tissue’, Innovative Food Science and Emerging Technologies, 3, 371–377. Zeuthen P and Bogh-Sorensen L (eds) (2000), Food Preservation Techniques, CRC Press, New York. Zhang Q, Monsalve-Gonzalez A, Qin B L, Barbosa-Canovas G V and Swanson B G (1994), ‘Inactivation of Saccharomyces cerevisiae in apple juice by square wave and exponential-decay pulsed electric fields’, Journal of Food Process Engineering, 17, 469–478. Zimmermann U (1986), ‘Electrical breakdown, electropermeabilization and electrofusion’, Rev Physiol Biochem Pharmacol, 105, 175–256.
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Advances in process chromatography and applications 109
4 Advances in process chromatography and applications in the food, beverage and nutraceutical industries M. Ottens and S. Chilamkurthi, Delft University of Technology, The Netherlands Abstract: The basics of process chromatography and its application to separation and purification in the food, beverage and nutraceutical industry are discussed. Basic principles and modes of chromatography are described together with guidelines for appropriate modeling; all illustrated with examples. New developments are described in the area of novel ligands, new materials, and modes of operation and control. Finally, future trends with respect to the production of high-value nutraceuticals, regulation and process control are given. Key words: process chromatography, nutraceuticals, proteins, model, simulated moving bed.
4.1 Introduction 4.1.1 Chromatography as a unit operation in food, beverage and nutraceutical processing The global food, beverage, and agriculture industry caters to the population of the entire world. It is a complex, global collective of diverse businesses that together supply much of the food energy consumed by the world population. Global market forces are driving the continuous evolution of the food and beverage industry. Consolidation, changing consumer preferences and increasing government regulations are significantly impacting manufacturing and business strategy. Producing quality goods at the lowest possible cost is a major aim. Consumers are increasingly interested in the health benefits of foods and have begun to look beyond the basic nutritional benefits of food to the
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110 Separation, extraction and concentration processes disease-prevention and health-enhancing compounds contained in many foods. This, combined with a more widespread understanding of how diet affects disease, health-care costs and aging populations, has created a market for functional foods and natural health products. There is a growing interest in functional foods, which are dietary supplements that increase a person’s well being. A functional food contains one or more ingredients that are bioactive, that is: they promote desired biological activity in the human body. Process-scale chromatography is an indispensable unit operation in biotechnology and finds wide application in the food and beverage industry (Fig. 4.1). Even though the food industry does not rely on process scale chromatography to the same extent as perhaps other industries such as the pharmaceutical sector, the sheer size of the industry results in a considerable application for chromatography. Other than process-scale applications of chromatography, many applications are related to safety, quality control, and research of new food ingredients and products. Many product streams in the food industry contain components that may have applications for, for example, functional foods. These product streams are typically large in volume and contain only a small amount of the component (or mixture of components) of interest. For economical application of large-scale preparative protein, peptide or oligosaccharide chromatography from agricultural/biotechnology streams, there is a need for ‘cheap’ resins (the costs for resins may be as high as 70% of total processing costs). The standard technology for fractionation or enrichment of food streams is using a packed bed. The close contact between interstitial liquid and carrier beads ensures effective mass transfer. For the treatment of large process streams
Fermentation Cell disruption Separation of cell debris
Removal of ‘like’ molecules such as mis-folds, multimers, degradation products via: Crystallization (extraction) Chromatography
Concentration Purification
Physical adsorption Ion exchange Affinity, size (sorption processes)
Formulation
(membrane separations) (electrokinetic separations)
Market
Fig. 4.1 Positioning of chromatography in a general layout of a bioprocess. For food processes, the fermentation step may not be present.
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Advances in process chromatography and applications 111 and/or process streams where suspended solids or fouling components are present, problems may arise: the pressure drop over the column might be too high. The workhorse of preparative protein purification is ion exchange chromatography (IEC). For dilute feeds, IEC can be used in a capture step. Another way of pre-concentration is precipitation followed by resolubilization and IEC fractionation. Large-scale preparative chromatography can be done more efficiently using counter-current chromatography, which has larger driving forces for mass transfer and more efficient use of desorbent and (expensive) resin (see Section 4.4). As an example of the fractionation of protein hydrolyzates, simulated moving bed (SMB) chromatography has been recently applied (Ottens et al. 2006b). An example of large-scale purification of biomolecules being applied in industry is, for example, the purification of lysine (an amino acid) in SMB carousels (Van Walsum and Thomson, 1997). However, using a larger number of protein fractions, more side streams are needed, or a different approach will be needed (similar to distillation with different side streams). Resin screening is an important aspect of the development of an economical process (with respect to resin capacity, regeneration cycles, resin lifetime, costs of resins, food grade resins: some resin is always lost in the feed and the need for ‘water-soluble elution liquids’, as many food companies are not used to processing organic solvents). Last but not least, fouling characteristics are important as well as a proper cleaning program. The current challenges that are associated with the use of chromatography on a large scale in the food industry are the following (see also section 4.4): Large diluted stream with minor component (mg m–3) from large product streams (m3 h–1). – For high throughputs, expanded bed adsorption (EBA) or SMB chromatography might be considered, as indicated above. Considerable effort in the application and development of EBA was made by Kula and co-workers (e.g. Hubbuch et al., 2006). ∑ Pressure drop. Possible solutions include: – use of micro-structured materials, multimodal pores, perfusion resins; monoliths, structured woven packing, all allowing for a large convective flow and creating an interfacial area in the (functionalized) micropores, – (partly) fluidization, as in EBA, – radial flow chromatography. ∑ Fouling: – removal of debris, proteins and fats from the feed by membrane filtrations preferably in combination with a pretreatment. ∑ An understanding of the required purity of the targeted molecule. Mixtures with enriched fractions or pure targeted component. ∑
These considerations are addressed in this chapter.
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112 Separation, extraction and concentration processes Several examples of functional food products that can be enriched or purified by chromatography are briefly introduced next. Peptides are known to have a versatility of bioactivity, e.g. ranging from a metal binding function, through influencing mood functions, to a reduction of blood pressure. The peptides can be obtained from various (food) proteins; dairy liquids such as milk and whey have often been investigated (e.g. Meisel and Schlimme, 1996). Typically, the production of a bioactive ingredient for functional food is obtained by enzymatic hydrolysis of the native protein. From the resulting complex mixture, the peptide(s) of interest need to be isolated. Currently, this is often done by membrane filtration or chromatography. Membrane filtration is regarded as a cheap technology, but the selectivity is rather low. Electromembrane filtration has been suggested as an alternative isolation route for charged components. In this membrane-based separation, the selectivity is increased by applying an electrical field as the (selective) driving force for migration over the membrane (Bargeman et al., 2002a; 2002b), which contributes to separation costs and complexity in comparison to regular membrane filtration. On the other hand, chromatography is highly selective, but is often regarded as expensive, as a result of the relatively high costs of the column media and eluents. The use of SMB technology is investigated as an alternative separation route, as this counter-current chromatographic separation method has been shown to be able to efficiently reduce sorbent inventory and the use of buffers by one order of magnitude in comparison to regular (fixed bed) column chromatography (e.g. Ballanec and Hottier, 1993; Cavoy et al., 1997). Furthermore, the use of chromatography for glucose–fructose separation (Azevedo and Rodrigues, 2005; Luz et al., 2008) sugar recovery from molasses (Hannu and Jarmo, 2000), stabilization of beers (Rehmanji et al., 2002), separation of proteins from various sources, such as milk (Kim et al., 2003a) and rapeseed, are only a few of the process scale applications of chromatography in the food and beverage industry. In addition to this several other functional components such as amino acids (Kostova and Bart, 2007a,b), flavors and fragrances (Gunther and Armin, 1993), nutraceuticals (Genovese et al., 2007), organic acids (Yoshikawa, et al., 2007), and oligosaccharides (Berensmeier and Buchholz, 2004) have also been separated on process and analytical scales. All these separations involve various interaction mechanisms, including ligand-exchange chromatography, LEC (Azevedo and Rodrigues, 2005, Luz et al., 2008, Steffenson and Westerlund, 1996), size-exclusion chromatography, SEC (Thurl et al., 1993), IEC (Kim et al., 2003a), and reversed-phase highperformance liquid chromatography, RP-HPLC (Stah et al., 1994; Thurl et al., 1991). The various interaction modes are described in the next section. Sanitation and hygienic design are key characteristics of food, beverage and nutraceutical processing. Major chromatographic companies that operate in the field are amongst others Universal Oil Products (UOP) and Novasep, and they can readily assist in designing large-scale chromatographic separations.
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Advances in process chromatography and applications 113 4.1.2 Main chapter themes In the remaining part of this chapter first the basic principles of chromatography are described and the various types and basic modeling outlined. Then some applications are presented to indicate how chromatography is currently employed and its future potential in the food field. Subsequently, a section on recent developments in the chromatographic field is given with implications and applications for use in the food industry. This section is followed by an overview of future trends in the food chromatography field and finally some conclusions and suggestions for further reading are given.
4.2 Basic principles of process chromatography 4.2.1 Types of chromatography used The different modes or types of chromatography that can be used are categorized and depicted in Fig. 4.2 (see also GE website for further information; GE Healthcare is one of the major suppliers of chromatography equipment and resins: http:\\www.gehealthcare.com). The various methods are ranked according to their interaction strength, with the weakest on top. Size-exclusion chromatography is based on differences in the size of the solute molecules, by using a porous resin with distinct pore sizes (Horneman et al., 2004, 2006). Hydrophobic interaction chromatography and reversed-phase chromatography involve separation based on the differences in (surface) hydrophobicity of the solutes. The most powerful and most often used in food processing is ion-exchange chromatography, which separates solutes based on the differences in charge, which can be modulated by the pH and ionic strength of the solution. The strongest and most selective method is Sorption mechanisms
Weak
∑ Size
∑ Gel filtration (GF)
∑ Van der Waals
∑ Hydrophobic interaction (HIC) and reversed phase (RPC)
∑ Polar
d+
d–
d+ d+ +
∑ Ionic Strong ∑ Affinity
+
+
∑ Ion exchange (IEC)
+ ∑ Affinity chromatography (AC)
Fig. 4.2 Illustration of the basic interaction modes of chromatographic separation.
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114 Separation, extraction and concentration processes affinity chromatography, and it is most often used in the pharmaceutical field, because of its selectivity and because its high cost can be met here. The purpose of process chromatography is to separate a target component from a mixture. Liquid chromatographic separations are based upon the different degrees of interaction of solutes dissolved in a mobile phase with a chromatographic medium or resin (the stationary phase), resulting in differential migration of the solutes through the column as illustrated in Fig. 4.3. The feed to the chromatography column is a binary mixture of solutes A and B, moving with concentration (wave) velocities wA and wB, respectively, through the column, with wB > wA. That is, solute B is more strongly retained or retarded by the resin than solute A. Each solute band becomes more dispersed as the chromatographic process progresses (depicted with fainter shading in Fig. 4.3b). Therefore, a chromatographic purification that is mainly based on adsorption and desorption of the target compound consists of several steps as follows. Equilibration of the resin For equilibration of the resin, the adsorbent material is prepared for the chromatographic separation. Depending on the mechanism of adsorption the equilibration step may include equilibration with the desired ionic strength, pH, and solvent concentration. Usually the equilibration solvent is chosen such that upon adsorption of the target compound and its impurities no change in composition (pH, ionic strength, and solvent composition) of the mobile phase occurs. In some instances it may be necessary to change the composition of the mobile phase to enhance the adsorption process. Loading In the loading process step, the target compound (and some impurities) is adsorbed onto the media. The capacity of the resin and/or the difference in affinity between target compound and its impurities determine the loading volume and loading conditions. It may be necessary to change or alter the feed stream properties to optimize the adsorption process. Washing In the wash phase, a washing buffer or solvent is used to desorb impurities from the resin. The washing buffer may be similar to the equilibration buffer, but this is not necessarily the case. Elution of the target compound needs to be avoided in this step. Elution In the elution phase the target compound is desorbed into an elution buffer solution. This can be done both in isocratic mode (constant mobile phase composition) and in gradient mode in which a gradient in the eluting buffer is applied to achieve optimal purity and/or minimal elution volume.
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Advances in process chromatography and applications 115 Feed Mobile liquid phase
WB
Mobile phase
WA
WB
Mobile phase
WA
WA
WB
More strongly retained Less strongly retained
Stationary solid phase
t = 0 min
t = 10 min
t = 20 min Concentration
tR,B-tR,A
A
(a)
B A
B A
tR,A
tR,B
wA
wB
Time
Regenerant (eluent)
Feed B A
B
Flush
A
A B
B
B
A
A
B
A
A
A
A
Time
(b)
Fig. 4.3 (a) Illustration of the basic principle of chromatographic separation. Chromatogram showing the exit concentrations of A and B as a function of time (arbitrary times are used to show the progress of the chromatographic separation). (b) Adsorption, flush and regeneration.
Sanitization After elution a sanitization step may be applied in which compounds that show a higher affinity to the resin than the target compound are desorbed. For polymeric media, often a solvent is used to remove very apolar compounds from the resin. For hydrophilic media, a caustic solution is often used to © Woodhead Publishing Limited, 2010
116 Separation, extraction and concentration processes hydrolyze any irreversibly bound proteins. Acidic washes are applied to remove salts with low solubility (often calcium salts). Varying the salt concentration, pH or solvent composition can also be used to deliberately shrink and swell the resin which may enhance sanitization. After sanitization, the process starts again from equilibration. For processes in which there is very little affinity difference between the target compound and a limited number of impurities, it is often decided to use a continuous mobile phase with constant composition. The compounds are collected sequentially at the exit of the column and no intermediate equilibration, washing and sanitization is applied. A typical example of such a process is the separation of glucose and fructose on strong cationexchange resins in the Ca form. The requirements for sorbents to be used in chromatography and the different basic types are outlined in Fig. 4.4. 4.2.2 Modes of operation There are three different modes of operating a chromatographic process: frontal, elution and displacement. These stem from the three different ways of desorbing bound solutes from the chromatographic medium. Requirements ∑ Insoluble
∑ Hydrophilic Compromise
∑ Macroporous ∑ Mechanically stable
∑ Proper shape ∑ (Bio-)chemical stability
Main types
Composite
Organic ‘polymers’
Inorganic ‘pellets’
(a) ∑ Inorganic
∑ Organic polymers
Activated carbon
Polystyrene
Silica
Poly(meth)acryl
Alumina
Polysaccharides (dextran, agarose, etc.)
Pellet
Dense gel
Apolar
Macroreticular
Polar
Open gel
(b)
Fig. 4.4 (a) Requirements for sorbents and the main types available. (b) Examples of inorganic and organic polymers used as sorbents.
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Advances in process chromatography and applications 117 In frontal chromatography, the sample is also used for elution. In this case, the sample feed is continuously applied to the column and the sample components displace each other in order of decreasing affinity for the chromatographic medium. The least retained solute is obtained in a pure form (i.e. depleted of the more strongly retained components) until the other solutes break through. Eventually, the column is saturated with the strongest retained component and the effluent has the same composition as the feed. In elution chromatography, the solutes are desorbed from the chromatographic medium by the action of a competing or modifying agent in the mobile phase, such as a salt. The concentration of the modifying agent is a key parameter. Three situations can be distinguished: isocratic, gradient or step change in the mobile phase composition. By keeping the composition (e.g. ionic strength) of the eluting buffer constant, the capacity factor is also kept constant; this is called isocratic elution and it is primarily used for separation of small molecules for which the variation of retention factor with mobile phase composition is not as large as for larger molecules. In gradient elution, the composition of the mobile phase is continuously changed (e.g. increasing ionic strength as in ion exchange chromatography), resulting in a continuous change in the retention factor. Instead of using a continuous gradient, the mobile phase composition may be changed in a stepwise fashion, resulting in discontinuous changes or ‘shifts’ in the retention factor. Gradient or step elution may be necessary for separation of large biomolecules for which the binding strengths may be too strong for isocratic elution to be feasible. In displacement chromatography, the affinity of the competing agent for the chromatography media is much higher than that of the solute, such that the agent effectively displaces the solute. This takes place irrespective of the concentration of the displacing agent (or displacer), in contrast to elution chromatography. Displacement chromatography has the advantage of being able to concentrate samples. However, selecting an appropriate displacer may not be easy. Several important aspects of using chromatography in food separations are outlined in Fig. 4.5. 4.2.3 Basic modeling The use of mathematical modeling is helpful in the design of chromatographic separators to size the chromatographic piece of equipment and to determine its optimal operation conditions. There is a wealth of information available on modeling of chromatography in different levels of detail. Guiochon gives a good basic overview (see section 4.8), and some recent modeling efforts using the general rate model and the linear driving force model for mass transfer are given by Nilsson (2005) and Horneman et al. (2006), respectively. Although the general rate model is the more accurate one, a less computationally involved but still accurate one is the linear driving force
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118 Separation, extraction and concentration processes Eluent recycle ∑ Operating conditions and short-cut design Raffinate product Feed
∑ Complex phase behavior (azeotropic separations) ∑ Non-equilibrium fractionation ∑ Reactive separations
Extract product
Eluent make-up
∑ Multiple side/product streams ∑ Column dynamics
Sorbent recycle
Fig. 4.5 Schematic diagram of the operation of a large-scale chromatographic fractionator, such as a simulated moving bed, and the main factors for successful operation.
model, as shown in Fig. 4.6. It contains two partial differential equations describing the concentration of solutes in the bulk liquid and solid phase as function of time and axial position in the column. The equations are connected by a mass transfer term containing the isotherm, which gives the ratio of partitioning over the solid and liquid phase of a solute under certain environmental conditions. Several isotherms can be used, the most common being the Langmuir one, which gives a linear relation between the concentration in the solid and liquid phase at low concentrations and levels off at a constant saturated value at higher concentration at the maximum loading capacity of the specific resin used. Any appropriate PDE solver can be used (e.g. PDESol, Comsol) to solve the set of obtained partial differential equations. Commercial packages that are available on the market are for example ASPEN Chrom, gPROMS, and for flowsheeting SuperProDesign (Intelligen).
4.3 Applications of process chromatography in the food, beverage and nutraceutical industries Process scale chromatography has been effectively used in the food, beverage and nutraceutical industry (Fig. 4.7 and Table 4.1). Examples of these applications are presented here and by doing so a brief overview is provided of the industrial areas where process chromatography is applied. The various sources and products that can be concentrated, fractionated and purified by chromatography are given in Fig. 4.7 and Table 4.1, as well as scales of some chromatographic separations actually in industrial practice with its operation mode (SMB, and fixed bed chromatography, FB).
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Advances in process chromatography and applications 119 4.3.1 Glucose–fructose separation The technology used previously (up to the middle 1970s) in the preparation of fructose was based on the application of a scarce and expensive raw material, inulin, which is a fructose polymer. This prevented the wide-scale use of fructose in the food and medical industry. Only in the 1980s did the application of chromatographic methods permit the large-scale manufacture of fructose by its isolation from glucose–fructose syrups obtained either as ∑ ∑ ∑ ∑
Which phases? Which convective streams? Is there dispersion? Mass transfer between phases?
Void fraction f
(Interstitial) velocity v
Feed flow Q MT
Cross-section A
Q = v fA
Mass transfer
Dispersion
(a) Mass balance over a slice Convection
(fA)vc
Dispersion
Ê ˆ (fA)Á – E ∂c ˜ Ë ∂z ¯
Mass transfer Accumulation
z + Dz
– [(1 – f)A] MTDz (fADz ) ∂c ∂t
Change = in – out ∂c = – ∂t Equation 1
z
Per volume solid
Dz (vc ) |z+ z
Dz
+E
Ê ∂c ˆ ÁË ∂z ˜¯
∂c = – v ∂c + E ∂2c – 1 – f MT ∂t ∂z f ∂z 2
z+Dz z
Dz
–
1–f MT f
E = 2vdp
(b)
Fig. 4.6 Mathematical modeling for chromatography (a). Different (partial differential) equations to be used (b) and (c), with ‘closing relations’ for mass transfer (d). Symbols: z axial length scale, D difference, c concentration, t time, k mass transfer coefficient, a interfacial area, K partition coefficient, Sh Sherwood number, Re Reynolds number, Sc Schmidt number, D diffusion coefficient, d particle diameter, n kinematic viscosity, v velocity, A cross- sectional area, Q flow rate, f void fraction, MT mass transfer, E dispersion coefficient. Subscripts: f film, p particle. Superscripts: * at equilibrium, vinculum solid phase.
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120 Separation, extraction and concentration processes z + Dz
z
Mass balance over a slice of solid Mass transfer MTDz = ka (c* – c) Dz Overall coefficient Area per volume of solids Equilibrium concentration Accumulation
∂c Dz ∂t
Equation 2
∂c = ka (c* – c) = MT ∂t
Overall MT coefficient k–1 = (Kkf)–1 + k–1 p
Solve two coupled PDEs (c) dp = 10–3 m
In the ‘film’ Shf = 4 + 1.1Re0.6 Scf0.33 kf dp
Shf ∫ Re ∫
Df
; Scf ∫
nf ; Df
vdp nf
Df = 10–9 m2 s–1 v = 10–3 m s–1 Re = 0.5 Scf = 1000 Shf = 10 kf = 10
Df dp
= 10–5 m s–1
In the ‘particle’ Dp = 10–10 m2 s–1
Shp = 10 Shp ∫
k pd p Dp
; Dp π Df
kp = 10
Dp dp
= 10–6 m s–1
(d)
Fig. 4.6 Continued
a result of the inversion of sucrose or by converting glucose with the aid of glucose isomerase. The most common type of fructose syrup, usually called high-fructose syrup, contains 42% fructose, 52% glucose, and 6% oligosaccharides on a dry basis, and is used in place of sucrose in some foods and beverages (Kishihara et al., 1992). For some purposes, because fructose is sweeter and more soluble in water at low temperatures than glucose, syrup with 55 to 90% fructose, called higher-fructose syrup, is desirable. Therefore, to produce syrup containing more than 50% fructose a process to separate fructose from the equilibrium mixture is desirable. This mixture can be separated by chemical and physical means. The most widely used method for this kind of separation is chromatographic separation using cation exchange resins in Ca2+ form (Barker et al., 1984). In this instance, SMB is a widely used technology (see Section 4.4.3) (Bubnik et al., 2002). © Woodhead Publishing Limited, 2010
Advances in process chromatography and applications 121
Dairy industry
Marine
Milk Eggs
Fish Algae Peptides Proteins
Omega-3, -6 fatty acids
Agriculture Sugars Flavors
Fig. 4.7 Source and product areas where process chromatography is being used or has potential. Table 4.1 Industrial scale chromatographic separations. Taken from Lecture notes from the Advanced Course on Downstream Processing, Delft, 2009, and Ganetsos and Barker (1992) (see Section 4.8) System
Process*
Scale (tonne a–1)
Xylene isomers Glucose–fructose Lactic acid, citric acid Amino acids Chiral compounds Peptides Milk proteins Flavors
SMB SMB FB/SMB FB/SMB SMB SMB FB FB
400 000 100 000 10 000 1500 1–10 0.1–1 0.1 0.01–01
*FB, fixed bed; SMB, simulated moving bed.
When large-scale separations are required, the SMB process is often the technology of choice in order to reduce separation costs. Compared with batch chromatography, SMB technology uses less eluent; in addition the separation productivity is higher thus using less adsorbent (Broughton, 1984; McCulloch et al., 1994). In recent years, fructose and glucose were separated from complex mixtures such as cashew apple juice using SMB technology (Azevedo and Rodrigues, 2005; Luz et al., 2008). 4.3.2 Recovery of sucrose from molasses Chromatographic separations are used to separate valuable components from beet molasses in beverage industries. Characteristic products that are © Woodhead Publishing Limited, 2010
122 Separation, extraction and concentration processes currently produced in addition to sugar include betaine, inositol, peptides, amino acid mixtures and several individual amino acids. The recovery of the other components can significantly improve the economy of sucrose recovery, but still sucrose remains as the most important product controlling the overall economy of the process. Chromatographic separation can also be applied to treat intermediate juices in the beet sugar industry. Thick juice separation has also been suggested. The resolution of the peaks may change owing to the changing composition of the feed material. Industrial batch systems were first built in the 1960s and the 1970s. In the mid-1980s the continuous simulated moving bed process was applied in molasses separation. The sequential SMB process, which can simultaneously recover multiple product fractions, was introduced in the late 1990s (Hyöky, 1999; Lancrenon, 1997; Paananen, 1996; Rousset, 1997). Hyöky et al. (1998) described the multi-profile FAST separation, which effectively doubled the efficiency. Further development includes a patented two-stage process (Hyöky et al., 1998). It is a combination of two chromatographic fractionators to improve the recovery and the purity of overlapping components, such as sucrose and betaine. It is uniquely suitable for recovering multiple value-added products from the process streams of the sugar industry (Paananen and Kuisma, 1999). Chromatographic systems are used in production of sweeteners based on beet, cane and starch as described by Paillat (1999). Sugar beet produces sucrose, betaine, raffinose, invert and many other water-soluble components. These components are extracted from the beets as juice. Part of these watersoluble components are removed or destroyed in the juice purification, but a large part of them will end up in the by-product: molasses. Molasses is the most common raw material in chromatographic desugarization plants. The main product is sucrose, but in addition betaine, inositol, amino acid mixtures and individual amino acids have been commercially separated from molasses (Hyöky et al., 1998; Paananen, 1996; Paananen and Kuisma, 1999). Low green or B molasses separation is already in industrial use. Kearney has suggested thick juice separation (Kearney, 1997; Kearney and Rearick, 1995). The positioning of the chromatographic step in the beet process has distinct effects on the recovery of sucrose and especially on the recovery of the other components. 4.3.3 Stabilization of beers Brewery production of some corns and hops for beer is a universal process in the beer industry. The quality of beer depends on its taste, clearness, color and foam retention. One of the key processes is the procedure of removing non-micro-organisms impurities. Fixed-bed chromatography with several resins like silica gel (SiO2nH2O) are in use to remove proteins that make beer muddy. Beers are known to contain a wide variety of phenolic compounds, most of which originate from the raw materials of brewing, i.e., barley and hops
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Advances in process chromatography and applications 123 (Gardner and McGuinness, 1977). Of these phenolic compounds, the flavanoids are of particular interest to brewers, as they have long been thought to be precursors of non-biological haze in unstabilized beers (McMurrough et al., 1983). The flavanols have been classified into three groups on the basis of their chromatographic behavior. The first group, the simple flavanols, comprise flavanol monomers [e.g. (+)-catechin and (–)-epicatechin)], dimers (e.g. prodelphinidin B3 and procyanidin B3), and trimers. The second groups, the polymeric flavanols, are formed by oxidation and polymerization of simple flavanols. The complexed flavanols, the third group, result from the interaction of polyphenols with proteins to form complex structures (McMurrough, 1979). Stabilization of beer against haze formation may be achieved by decreasing the simple flavanol content, thereby limiting further flavanol polymerization and complexation. Increasingly, this stabilization is being achieved in breweries by treatment of beer with polyvinylpolypyrrolidone (PVPP) before packaging (McMurrough, 1993; McMurrough et al., 1992). Stabilization is achieved by sorption of phenolics on PVPP, which is subsequently removed by sheet filtration. This procedure can be carried out in conjunction with silica hydrogel treatment to remove haze-forming proteins (McMurrough and Madigan, 1996). These amorphous, non-additive stabilizers offer selective removal of haze forming proteins, improving colloidal stability of beer with no adverse effects on head or taste. Typically these silicas have been used with body feed diatomaceous earth filter aid (D.E.). However, by changing the conditions in the manufacturing process, it is possible to vary permeability (filterability) of silica hydrogel products. An innovation in the stabilization of beer is the ‘combined stabilization system’ (CSS) which is capable of a combined removal of turbidity forming protein and polyphenols in a single step (Janey and Katzke, 2002). As the economic alternative to PVPP and silica stabilization, the CSS is a compact rig mounted fully automated stabilization system, which can be integrated into any existing filter line. The adsorbent is permanently retained between an inlet and an outlet screen, making the dosing of precoat and feed suspension before each filtration unnecessary. A CSS adsorber has a particle size of 100– 300 mm. The solid phase adsorber resin is based on a high-grade, cross-linked, insoluble agarose (polysaccharide). Protein and polyphenols are adsorbed and then removed from the agarose adsorbent during the regeneration. Neither substance is dissolved in the beer nor is the beer quality affected in a negative way. The beer’s organoleptic properties, its foam stability, color and bitterness units remain unaltered with this kind of approach. 4.3.4 Separation of lysozyme from egg white Lysozyme has a number of possible applications; for example it can be used as an additive to baby milk or ophthalmic preparations; for treatment of ulcers, wounds and infections, as a potentiator of some antibiotics; as an
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124 Separation, extraction and concentration processes antioxidant; and for its antimicrobial properties (Owen and Chase, 1997). The expanding potential for application of lysozyme in many fields of science and technology dictates the development of efficient and simple methods for lysozyme purification. Several chromatographic techniques that have been used for the separation of hen egg white lysozyme include ion exchange chromatography (Banka et al., 1993; Li et al., 1999; McCreath et al., 1997), affinity chromatography (Chiang et al., 1993; Yamada et al., 1985; Yamasaki and Eto, 1981), dye-binding chromatography (Tejeda-Mansir et al., 2003), affinity membrane separation (Arica and Bayramoglu, 2005; Bayramoglu et al., 2003; Ruckenstein and Zeng, 1997), ultrafiltration (Ghosh and Cui, 2000; Ghosh et al., 2000), PEG/salt aqueous two-phase system (Su and Chiang, 2006), reverse micelles (Noh and Imm, 2005), metal-affinity precipitation (Roy et al., 2003) or adsorption to plant residues (Hou and Lin, 1997). Commercial cation-exchange resins such as Duolite C-464, Amberlite, CM Sephadex have been used for the separation of lysozyme on process scales with recoveries of 90–95% (Li-Chan et al., 2006). Nonetheless, the lengthy steps and dilution of product at the end of the process have hindered their application at process scales. More recently this problem has been overcome by the use of magnetic separation techniques for the separation of egg white lysozyme and several other proteins (Safarik and Safarikova, 1993; 2004). These techniques usually enable simple, onestep separation of target proteins, and in most instances magnetic affinity adsorbents can be used. However, large-scale application may be difficult to achieve. 4.3.5 Separation of whey protein from milk Milk proteins are the most important source of bioactive peptides. Typical bovine milk contains 13% solids, with 4% fat present as an emulsion of globules with diameters up to 10 mm and caseins present as a colloidal suspension of particles with diameters up to 0.1 mm (Bylund, 2003). Fat globules normally cause problems for chromatographic separations, as they block packed columns as soon as the feed is introduced. Raw whole milk contains larger suspended particles than whole (full fat) processed milk, as the latter is homogenized to produce a uniform consistency. The composition of bovine milk includes water, fat, lactose and minerals, and up to 6% of the mass is made up by proteins and peptides, among them a number of high-value substances (Lourdes et al., 2000). In particular, milk contains two major protein groups, caseins and whey proteins, which differ greatly with regard to their physicochemical and biological properties. Normal milk contains 30–35 g L–1 proteins, approximately 80% of which are caseins with the remainder being the whey proteins (Korhonen et al., 1998). Whey proteins can be acquired as a by-product in cheese manufacturing process. The required long-term stability in functional performance of these proteins is often lacking. In general, whey is dilute liquid composed of lactose, a variety
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Advances in process chromatography and applications 125 of proteins, minerals, vitamins and fat. Whey contains about 6% solids of which 70% or more is lactose and about 0.7% is proteins (Gerberding and Byers, 1998). Whey protein components are a-lactalbumin, b-lactoglobulin, immunoglobulins A, M and G, bovine serum albumin (BSA), lactoferrin and lactoperoxidase. b-Lactoglobulin is the major whey protein in bovine milk; it has a molecular weight of 18.4 kDa, possesses 162 amino acid residues and its concentration is 2–4 g L–1. a-Lactalbumin is an albumin which has 123 amino acid residues. It possesses a molecular weight of 14.2 kDa and its concentration in milk is 0.6–1.7 g L–1 (Ye et al., 2000). Bovine whey proteins have potential applications in veterinary medicine, food industry and as supplements for cell culture media. Immunoglobulin G (IgG) and immunoglobulin A (IgA), present in bovine whey, have high pharmaceutical value (Hahn et al., 1998). a-Lactalbumin can be used in infant formula and as a nutraceutical because of its high tryptophan content. b-Lactoglobulin is used in the production of confections (Zydney, 1998). Oral administration of bovine IgG is known to be an effective treatment of various infections of newborn infants (Hutchens et al., 1990). Typical concentrations, molecular weights and isoelectric points of whey proteins are given in Table 4.2 (Al-Mashikhi et al., 1988). Several techniques are used for the partitioning of whey proteins by e.g. aqueous two-phase systems (Jose et al., 2000; da Silva and Meirelles, 2000; Rito-Palomares and Miguel, 1998). Also, some attempts have been made to isolate whey proteins by using membrane filtration (Lucas et al., 1998; Gerd et al., 2000). The separation of whey proteins by using ion exchange chromatography on a process scale has been investigated by many researchers and several methods have been reported (Gerberding and Byers, 1998); Ionexchange separations take advantage of electrostatic interaction between surface charges on biomolecules, such as amino acids or proteins, and clusters of charged groups on the resin phase. An adsorbing biomolecule displaces counterions associated with the surface, discharging a complementary buffer salt in the process. Adequate buffering is required to shield native protein structures from changes in pH adjacent to exchange surfaces (Donnan effect) and pH effects were induced by sorption. Selection of an appropriate buffer Table 4.2 Typical concentration of whey proteins and their isoelectric points (pI) Protein
Approximate MW (kD) concentration (w/w%)
Isoelectric point (–)
b-Lactoglobulin a-Lactalbumin Immunoglobulins (A, M, G) BSA Protease-peptones Lactoferrin (LF) Lactoperoxidase (LP)
0.3 0.07 0.06 0.03 0.14 0.003 0.002
5.35–5.49 4.2–4.5 5.5–8.3 5.13 3.3–3.7 7.8–8.0 9.2–9.9
18.4 14.2 150–900 69 37–55 78 78
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126 Separation, extraction and concentration processes is critical to the success of ion-exchange separations (Keith and Edwin, 1995). 4.3.6 Separation of lactoferrin and lactoperoxidase In addition to the caseins and whey proteins, which are generally isolated from milk, some specialty proteins are also identified as byproducts in these separations. Two very important such proteins of commercial interest, which are part of the whey protein fraction, are lactoferrin (LF) and lactoperoxidase (LP). Although production of such high-value whey proteins is a commercial reality, two aspects of dairy processing may not be optimal for their production. First, the proteins are subjected to a series of processing steps before being extracted. It is a generally accepted principle of bioseparation process design that proteins should be separated from a source material as fast and in as few steps as possible to avoid loss of activity and yield (Ladisch, 2001; Harrison et al., 2003; Ottens et al., 2006a). Lactoferrin (LF) and lactoperoxidase (LP) being basic proteins are usually captured from whey or skim milk by cationexchange chromatography and sold as specialty ingredients (Lonnerdal and Carlsson, 1977). The costs associated with the separation of these types of proteins is high, but the value and wide range applications in various functional food and nutraceutical products outweigh the cost. Lactoferrin, a 70–80 kDa single subunit glycoprotein, consist of two domains, each with one iron-binding site that requires synergistic binding of a bicarbonate anion (Al-Mashikhi and Li-Chan, 1988). A conformational change in the binding site, between open or closed forms, accompanies iron binding and release. The affinity of lactoferrin for iron is much greater than of transferrin which functions as an iron carrier in serum. It is present in human milk, and many exocrine fluids (saliva, mucous secretion) as well as in mammalian. Lactoferrin is a complex molecule with a number of interesting functional properties. Based upon these, lactoferrin may enter the composition in a wide range of products. Lactoperoxidase belongs to the category of peroxidase and acts as a biopreservative. Both these products find several applications based on their antimicrobial and antiviral activities such as in baby foods, food supplements, cosmetics, oral care, and preservation of meat. The production of high-value dairy proteins such as lactoferrin and lactoperoxidase normally requires extensive pre-treatments of milk to remove fat and caseins by centrifugation, precipitation, Ca2+ chelation and/ or filtration. Similarly, fat and caseins are normally removed before capture of recombinant proteins from the milk of transgenic animals (Morison and Joyce, 2005). Such pre-treatments can result in significant loss of protein yield and/or activity. Currently, high-value dairy proteins are viewed as by-products, with the major income of the industry coming from commodity dairy foods such as milk powder, cheese and butter. Economies of scale for production of commodity
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Advances in process chromatography and applications 127 dairy products mean that centralized processing is the industry norm. Milk is typically cooled and held at 4 °C in vessels on the farm for up to 2 days before being transported to a dairy factory. There, it is pumped to holding tanks and then undergoes a series of unit operations such as cream (fat) separation, pasteurization, homogenization and blending for standardization before further processing into individual products (Bylund, 2003). After a number of such operations (which varies from factory to factory), LF and LP may be extracted from skim (low fat) milk or, more commonly, from whey, which is produced as permeate during membrane concentration of milk or after precipitation of caseins much further down the processing chain as a by-product of casein production or cheese making (Tomita et al., 2002). Extensive pre-treatments of milk and whey before ion-exchange capture of proteins are by no means restricted to industry but are also used in the laboratory. Many authors have examined the capture and analysis of whey proteins by chromatography (Al-Mashikhi and Li-Chan, 1988; Al-Mashikhi and Nakai, 1987; Andrews et al., 1985; Chaplin, 1986; de Frutos et al., 1992; Donnelly, 1991; Doultani et al., 2004; Elgar et al., 2000; Etzel et al., 2000; Felipe and Law, 1997; Francis and Regester, 1995; Geberding and Byers, 1998; Hahn et al., 1998; Humphrey and Newsome, 1984; Konecny et al., 1994; Lonnerdal and Carlsson, 1997; Morr and Ha, 1993; Noppe et al., 1999; Torre and Cohen, 1996; Visser et al., 1991; Xu et al., 2000; Ye et al., 2000; Yoshida, 1988). Hahn et al. (1998) examined the performance of a range of commercially available pharmaceutical grade cation exchangers for protein capture from acid whey. Doultani et al. (2004) used cationexchange chromatography to produce a number of protein products from mozzarella cheese whey. Ye et al. (2000) used both anion and cation exchange chromatography to isolate a-lactalbumin, b-lactoglobulin, lactoferrin and lactoperoxidase from rennet whey. Cation exchange membrane chromatographic systems have been successfully used for the separation of lactoferrin and lactoperoxidase on an industrial scale from sweet whey concentrate with yields of more than 90% for lactoferrin. It has also been demonstrated that these system can be scaled up to handle 1 ¥ 105 tonne/year of whey (Plate et al., 2006). 4.3.7 Separation of napin from rapeseed meal Preparative chromatography of napin and other proteins from rapeseeds was successfully carried out with a combination of ion exchange followed by hydrophobic interaction and size-exclusion chromatography. The yield of this combined process was 18% for napin and 40% for cruciferin (Berot et al., 2005). Rapeseed proteins were underexploited as it was being used only as animal feed until its non-food applications were demonstrated in a European project (Green chemicals and biopolymers from rapeseed meal with enhanced end-user performance. European contract Enhance QLK5 CT 199901442). Hence large quantities were needed to demonstrate its potential © Woodhead Publishing Limited, 2010
128 Separation, extraction and concentration processes in functional additives. Preparative purification processes of seed storage proteins were demonstrated for 11S and 7S globulin-type proteins in pea and soy (Crevieu et al., 1996; Gueguen et al., 1984; Howard et al., 1983; Lehnhardt et al., 1983; Larre and Gueguen, 1986; Nielsen, 1985; Thanh and Shibasaki, 1976). But with rapeseed storage proteins, which are composed of 11S globulin and 2S albumin-type proteins, only analytical purifications were developed, except procedures based on ammonium sulfate selective precipitation of proteins (Raab and Schwenke, 1984). SEC was successfully used to achieve the separation of 11S and 2S proteins (Dalgalarrondo et al., 1986), and cation-exchange chromatography (CEC) to separate 2S isoforms (Monsalve and Rodriguez, 1990). More recently a new procedure for fractional enrichment and purification of plant (e.g. soy beans) and milk-based proteins was demonstrated by the use of volatile electrolytes for isoelectric precipitations (Hofland et al., 2003; Golubovic et al., 2005). Rapeseed protein meal has two major classes of seed storage proteins: 12S globulin (cruciferin), which represents 25–65% of its protein content (Raab et al., 1992) and 2S albumin (napin). In addition to these two proteins it also contains some minor proteins of interest, such as thionins, trypsin inhibitors and a lipid transfer protein (LTP). From the structure and physicochemical properties of these proteins it is evident that cruciferin shows very different characteristics from other proteins including: high molecular weight, neutral isoelectric point pI; but aggregation could occur during the purification process. On the contrary, napin and LTP show rather close characteristics in their molecular weights and pIs, which could complicate the purification process. Separation of napin and cruciferin can be obtained by a combination of nanofiltration in combination with several chromatographic steps of cationexchange, size-exclusion and hydrophobic interaction chromatography on process scales.
4.4 Recent developments in process chromatography Figure 4.8 shows some of the areas where there is potential for development in process scale chromatography. The innovations can be categorized in several fields, materials, modes of operation, and equipment. In the subsequent subsections some of the innovation in these areas will be highlighted. 4.4.1 Structured matrices and monoliths Monoliths for enhancing reactions have been around for several decades. For example, PolyHipe is a highly porous macrocellular cross-linked styrene–divinylbenzene (DVB) copolymer, prepared by polymerization of a high internal phase emulsion (HIPE) of water droplets dispersed in a styrene/DVB continuous phase (Ottens et al., 2000). This type of monolith © Woodhead Publishing Limited, 2010
Advances in process chromatography and applications 129 ∑ Materials – optimization of selectivity, MT, and hydrodynamics, magnetic beads, monoliths, structured matrices ∑ Mode of operation – exploiting non-linear effects of coupled multi-component equilibria ∑
Equipment – CPC: centrifugal partitioning chromatography – Radial flow: annular chromatography – SMB (counter-current chromatography) – EBA (expanded bed adsorption)
∑ Chromatographic reactors – coupled reaction – separation
I Feed A,B
Waste
II III IV
Solids B Raffinate A Extract
Desorbent
Fig. 4.8 Areas of innovation in chromatography and a schematic diagram of a simulated moving bed.
was used to enhance a multiphase reaction, but had structural properties that also made it suitable for high capacity, low-pressure drop chromatographic separations. Structured monoliths from Corning that have been used in chemical conversions in the work of Kreuzer et al. (2005) would also suit chromatographic separations purposes. Pictures of monoliths are given in Fig. 4.9. For more than a century, chromatography practitioners have been separating the components of chemical mixtures by using columns packed with various types of particulate matter. Monoliths are a relatively new class of stationary phases, completely different from conventional stationary phases. A good overview of this relatively new stationary chromatographic phase in the biotechnology area is given by Jungbauer and Hahn (2004). The material is cast into a chromatography column as a continuous block of one piece interlaced with channels (Jungbauer et al., 2002a). The ramified channels do not have dead ends. Owing to this structure, the transport of the solute to the surface is solely by convection instead of diffusion as observed in conventional media (Hahn and Jungbauer, 2000; Hahn et al., 2002;). These nontraditional column materials have recently been commercialized and are being recommended by manufacturers and users for the enhanced speed and thoroughness with which they can separate complex mixtures of biological molecules. The most important characteristics of monolithic media are the excellent mass transfer properties and the low pressure drops. The large channel diameter makes monoliths excellent stationary phase materials for chromatographic separations. The first monoliths were developed by Hjerten et al. (Hjerten and Liao, 1988; Hjerten et al., 1989, 1992) and Tennikova et al. (1990). Hjerten and coworkers compressed polyacrylamide gels and observed excellent resolution. Scale up was initially difficult but recently was successfully performed (Podgornik et al., 2000). Another important development is in the use of monoliths as a support for solid-phase synthesis (Jungbauer et al., 2002c). An interesting application is the direct synthesis of peptides on © Woodhead Publishing Limited, 2010
130 Separation, extraction and concentration processes
(a)
(b)
Fig. 4.9 Photographs of monoliths: (a) SEM photograph (15 kV; scale, 1 cm = 32.7 mm) of PolyHipe-type X20PV90 on an ISI-DS-130 SEM apparatus (Ottens et al., 2000); (b) Corning monoliths, as used in catalytic multiphase conversions by Kreutzer et al. (2005), with potential for use in large-scale food, beverage and nutraceutical purification.
monolithic columns. Because the synthesis is performed on polymethacrylate monoliths, the directly grown peptide can be used as an affinity ligand without any further treatment (Pflegerl et al., 2002a; 2002b; 2002c). This strategy provides an excellent screening platform for affinity ligands. The screening of the ligand can be performed by microtitration and the same resin can be used for screening and large-scale separation. Monoliths are also interesting supports for enzyme reactors. In packed-bed reactors, the efficiency is often limited by pore diffusion. The high porosity of monoliths ensures that the enzymatic process is not mass transfer limited (Jungbauer et al., 2002b). © Woodhead Publishing Limited, 2010
Advances in process chromatography and applications 131 A relatively new stationary phase is the combination of chromatographic resin beads and a membrane. Mixed matrix membranes (MMMs), which incorporate adsorptive particles during membrane casting, can be prepared simply and have performances that are competitive with other membrane chromatography materials. The application of MMM chromatography for fractionation of b-lactoglobulin from bovine whey was recently shown (Saufi and Fee, 2009). The dynamic binding capacity of b-lactoglobulin in whey solution was about 80 mg g–1 membrane (24 mg mL–1 of MMM), which is promising for whey fractionation using this technology. 4.4.2 Novel ligands Affinity separations are normally very expensive; therefore, application in the food industry requires either very high value products or a lower ligand cost. Affinity ligands are the basic moieties on which the various affinity chromatographic separations run. Affinity chromatography has been used widely in biomedical research and biotechnology (Wilchek and Chaiken, 2000). It is based on molecular recognition where one recognition partner is immobilized on a base matrix, and soluble target molecules can be retained from a crude mixture. The target molecule can then be released and recovered in a functional form. The basis of elution is to reduce the affinity between immobilized ligand and analyte. This is most often accomplished by use of a bond-breaking buffer (i.e., by changing pH, ionic strength, and solvent) or by use of competitive elution (Firer, 2001). In either instance, the diffusion of immobilized ligand is always strictly limited, and therefore, the local concentration of immobilized ligand is constant. The environment surrounding the immobilized ligand changes to modulate the ligand–analyte affinity. Affinity chromatography is potentially the most selective method for protein purification. The technique has the purification power to eliminate steps, increase yields and thereby improve process economics. Although affinity chromatography is used extensively on a laboratory scale, its widespread acceptance has been limited on the preparative scale because of the high cost of the affinity ligands and their biological and chemical instability (Lowe et al., 2001). Only recently, the development of new methods for screening, selection and design of stable synthetic ligands, has opened the opportunity of exploitation of such materials on a large-scale (Ladner and Ley, 2001). The rapid growth of bioinformatics and molecular docking techniques and the introduction of combinatorial methods for systematic generation and screening of large numbers of novel compounds, has made feasible the rapid and efficient generation of ligands for affinity chromatography (Clonis, 1990; Labrou, 2002; 2003; Labrou et al., 2004; Spalding, 1991). Immobilized lectin affinity chromatography (Endo, 1996; Kobata, 1994) using Ricinus communis agglutinin (RCA) has been used for separation of several glycopeptides and oligosaccharides (Merkle and Cummings, 1987;
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132 Separation, extraction and concentration processes Shinohara et al., 1987). RCA binds specifically to nonreducing-end galactose residues (Wei and Koh, 1978). Competitive elution using lactose, a hapten sugar, is generally used to elute the target molecule from the RCA immobilized column. Separation of bioactive peptides from defatted sunflower meal using immobilized lung extracts or immobilized angiotensin converting enzyme (ACE) (Megıas et al., 2004; 2006a; 2006b; 2007a; 2007b) has been used. Peptides that have angiotensin converting enzyme (ACE)-inhibitory activity are of great interest because of their potential antihypertensive effect, and have been found in hydrolysates of plant and animal origin (Murray and FitzGerald, 2007; Ottens et al., 2006a; 2006b). These are a few examples of using novel ligands in the separation of several valuable products from food sources. Enrichment by affinity chromatography generally allows further purification using classical chromatographic techniques. 4.4.3 Mode of operation Counter-current chromatography: simulated moving bed In conventional chromatography, the resin is fixed in a column, and the liquid flows through the column. As in extraction and distillation processes, the counter-current movement of the two phases improves the efficiency of the separation process, by establishing a higher driving force for mass transfer. A typical moving bed system is shown in Fig. 4.10b. It is equivalent to a distillation column. The liquid flows from section I to section IV, the sorbent moves from section IV to section I. A feed mixture (F) is introduced in the middle of the system. It is fractionated in an extract stream (E) containing the more retained component(s), and a raffinate stream (R), containing the less retained component(s). Similar to distillation, the movement of the components depends on the flow rates of the liquid and solid phases and the equilibrium distribution coefficient over both phases. Commonly, a four section SMB is used, the middle sections II and III serving to achieve the actual separation and section I cleaning the sorbent for reuse while section IV cleans the liquid for (partial) reuse. A real continuous movement of the solid phase is difficult to achieve in chromatography without losing resolution, owing to attrition of the solid phase. In practice, the movement of the solid phase is ‘simulated’ by periodic switching of a carousel of fixed chromatographic columns (Fig. 4.10a). The efficient use of the resin and liquid (desorbent D) is an important contribution to the efficiency of the SMB compared with that of fixed bed (FB) chromatography. Normally a reduction by a factor of 2–10 in resin and buffer use per amount of product produced can be achieved in an SMB compared with a FB. Chiral separation was the main driving force for the establishment of SMB separations in the fine chemicals and pharmaceutical industry over the past two decades (e.g. Cavoy et al., 1997; Pais et al., 1997; Rodriques and Pais, 2004). It is now a well-known and regularly applied chiral binary separation
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Advances in process chromatography and applications 133 Feed Liquid flow cycle
Liquid flow direction
Raffinate
Extract Liquid flow direction
Desorbent
(a) Sorbent recycle q4
q3
q2
IV c4
c3
II c2 c¢2
R
q0
q1
III
cF F
I c1
c0
D
E
N species: 11 N unknowns (c, q)0, …, 4 + c¢2 11 N independent equations (all linear)
4N concentration ratio 4N overall mass balances 3N balances: feed, desorbent, sorbent [A] x = b (b)
Fig. 4.10 (a) Schematic of the concept and operation of continuous multi-column simulated moving bed chromatography. Different tints indicate the extent of concentration of the two solutes. (b) A simple steady-state model for mathematical modeling of four-zone linear SMB chromatography: q solute solid-phase concentration, c solute liquid-phase concentration, R the raffinate stream, F the feed stream, E the extract stream, D the desorbent stream. The different zones are indicated by I–IV and 1–4.
technique. SMB is also a well-studied operation. However, some areas of application need to be developed. Separations of biologicals (Houwing, 2003; Houwing et al., 2002, 2003a; 2003b; Paredes et al., 2005) as well as multi-component (MC) separations in SMBs (e.g. Abel et al., 2004; Wang and Ching, 2005) are important research fields with a large industrial relevance. Multi-component (MC) separations can be performed in SMBs and some studies are devoted to it, mainly by serial SMB systems. Multiplezone (or section) SMBs have been reported (up to nine zones, Wooley © Woodhead Publishing Limited, 2010
134 Separation, extraction and concentration processes et al., 1998). Five-zone SMBs are being investigated and design criteria have been presented, for example, by Kim et al. (2003b). Wang and Ching (2005) presented a straightforward extension of a four-section SMB to a five-section SMB. An extra raffinate or extract port was added to the standard four-section SMB to perform a ternary separation. The pros and cons of two raffinate or two extract streams were discussed. However, in all these cases, a single component was purified from a three-component mixture. An equal or larger volume of industrial separations, particularly in the food industry, concern enrichment of a (MC) product in a particular functionality, for instance from a health or nutritional point of view. When using chromatography, the functionality can be associated to a particular MC fraction of an (also) MC feedstock. Developing and optimizing an efficient separation has to deal with the inherent MC nature of the feedstock and product(s). Also in chromatographic separations in food and biotechnology, consumption of costly resin and salty buffers is high, which are reasons to use SMB technology for these applications as well. As yet, few biotechnology applications of SMB technology have been described in open literature. The few examples found mainly concern small biomolecules, such as amino acids (e.g. Maki et al., 1987; Van Walsum and Thompson, 1997; Wu et al., 1998) disaccharides (Geisser et al., 2005) polypeptides (e.g. Mun et al., 2003) but also protein separations and purifications (e.g. Adachi, 1994; Gottschlich et al., 1996; Gottschlich and Kasche, 1997; Hashimoto et al., 1988; Horneman et al., 2006, 2007a; Houwing, 2003; Houwing et al., 2002, 2003a; Huang et al., 1988), plasmid purifications (Paredes et al., 2005) and viral clearance (Horneman et al., 2007b). In biotechnology, a whole range of chromatographic techniques is available for separation of proteins. SMB is applicable to any of these systems. Considering size exclusion; several forms have been investigated: separation of dilute mixtures (i.e., the linear isotherm case), surfactant-aided gel filtration (Horneman et al., 2004, 2006, 2007a; 2007b), separation of concentrated mixtures (the non-linear isotherm case) (Houwing, 2003), and separation of multicomponent mixtures (Ottens et al., 2006a). In the latter work, the technological and economic feasibility of SMB technology for multi-component product (and feedstock) separations is described, using the fractionation of peptides from protein hydrolysates as functional food ingredients as an example. The impact on the environment was also assessed. The model system consisted of an enzymatic hydrolysate of casein, which was fractionated using gel filtration (size-exclusion chromatography). The technical feasibility of isolating peptides, capable of inhibiting the ‘angiotensinconverting enzyme’ (ACE), from a casein hydrolysate by means of size-exclusion chromatography was investigated experimentally. In another study, the successful separation of the disaccharide lactose from a complex mixture of human milk oligosaccharides (HMOS) with the continuous chromatography of SMB technique is described (Geisser et al., 2005).
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Advances in process chromatography and applications 135 Expanded bed adsorption A major concern when operating a chromatographic column is fouling and, clogging of the column causing high pressure drop and preferential flow, or channeling, and thereby loss of capacity. To overcome this problem a technique was developed over recent decades that overcomes this problem, by operating the adsorption and washing step in the fluidized mode. Liquid enters the column from below at sufficient velocity to drag the sedimented solid stationary material with it, thus creating a fluidized or expanded bed. This mode of operation allows processing of crude feed streams with a high load of particulates, and can be used as a combined solid–liquid separation step and an initial capture step (e.g. Hubbuch et al., 2006). A recent application of expanded bed adsorption (EBA) in the food industry is the recovery fractionation and purification of valuable proteins from potato juice (Løkra et al., 2009). Horneman et al. (2008) used EBA in SMB mode to optimally recover and purify potato proteins to obtain an economically viable process.
4.5 Process control in chromatography In order to have an optimal separation at high productivity, yield and purity, it is necessary to operate under the appropriate conditions. The conditions may change owing to fouling or loss of capacity, a fluctuating feed composition or the inherent unsteady state of the operational procedure (i.e. with SMB chromatography). To operate at optimal conditions, necessitates control of the operation of the industrial-scale chromatography column(s). For SMB a safe regime is usually 10% below the optimal boundary. Proper control strategies may allow higher productivity, while maintaining the production constraints of purity and yield (Grossmann et al., 2008). With the availability of massive, cheap computer power and detailed mathematical models (as outlined in section 4.2.3), process scale chromatography will benefit more and more in the near future from proper control of the operation of the unit.
4.6 Future trends 4.6.1 High-value nutraceuticals Functional foods and nutraceuticals provide an opportunity to improve health while reducing health care costs. Functional foods involve the broad class of prebiotics and probiotics. Probiotics are viable microbial dietary supplements that influence the host and have a beneficial effect in the gastro-intestinal tract (Jardine, 2009; Salminen et al., 1998), whereas prebiotics are nondigestible food ingredients that benefit the host organism by stimulating the growth or activity of one or limited number of bacteria in the colon (Gibson and Roberfroid, 1995; Jardine, 2009). © Woodhead Publishing Limited, 2010
136 Separation, extraction and concentration processes Nutraceuticals refers to extracts of foods claimed to have a medicinal effect on human health. An important trend in food technology is the production of these ingredients from sources such as aromatic plants or spices. Among the nutraceuticals, antioxidants receive much attention in the food industry (Madhavi et al., 1996), not only as preservatives in food products to prevent or retard oxidation of fats and oils, but also because of their beneficial effects on human health. Supercritical-fluid chromatography (SFC) is one of the widely used process options in nutraceutical separations. Preparative-scale SFC is an environmentally clean technology whose main advantage, compared with preparative LC, is the easy recovery of the isolated compounds by a simple decompression of the supercritical fluid (Coleman et al., 1999). Fish oil preparation and separation of polyunsaturated fatty acid esters such as docosahexaenoic acid (DHA) and eicosapentaenoic acid (EPA) (Alkio et al., 2000) is a well-documented example of preparative SFC and has been used in nutraceutical production since the early 1990s. Novel methods such as continuous membrane chromatography reactor system (CMCRS) are used for the synthesis and separation of galacto oligosaccharides (GOS) (Engela et al., 2008). Membrane separations using ion-exchange chromatography is another widely used technique for functional components from milk sources (Goodall et al., 2008). 4.6.2 Good manufacturing practice (GMP), quality assurance (QA) and quality control (QC) and Food and Drug Administration (FDA) regulation Good manufacturing practice (GMP) is a term that is recognized worldwide for the control and management of manufacturing and quality control of foods, pharmaceutical products, and medical devices. GMPs are guidelines that outline the aspects of production that affect the quality of a product. The guiding principle of GMP is that quality is built into a product, and not just tested into a product. Therefore, the assurance is that the product not only meets the final specifications, but that it has been made by the same procedures, called standard operating procedures (SOP), under the same conditions each and every time it is made. In the food industry GMPs address all factors involved in a manufacturing process: personnel, building, premises, machinery and apparatus, documentation, and quality control, that generally affect product quality or impact quality monitoring. Besides these factors, an important theme that is part of GMP is validation. It is that part of GMP that ensures that facility systems, equipment, processes, and test procedures are under control and therefore consistently produce quality product. This is a special quality assurance method, which is explicitly required by GMP codes and mandated by law. GMPs are enforced in the United States by the FDA (http://www.fda.gov); within the European Union, GMP inspections are performed by National
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Advances in process chromatography and applications 137 Regulatory Agencies, e.g. GMP inspections are performed in the United Kingdom by the Medicines and Healthcare products Regulatory Agency (MHRA)(http://www.mhra.gov.uk), in the Republic of Korea (South Korea) by the Korea Food and Drug Administration (KFDA), and by similar national organizations worldwide. 4.6.3 Process analytical techniques (PAT) and control Process analytical techniques (PAT) are now being implemented in the pharmaceutical industry to monitor and control proper operation of these processes. Also in chromatographic fractionating and purification in the food industry such an application is seen. It allows for GMP, QC, better operation, better control, higher productivity and on spec production/higher number of on spec batches.
4.7 Conclusions Process chromatography is being applied in food, beverage and nutraceutical processing. Chromatography has a high resolving power which can give highpurity food products. With the advent of the nutraceutical field, also more expensive forms of chromatography, based on affinity, may find application, as the potential higher sales revenues balance the higher operating costs. Chromatography is a flexible purification method that is likely to play a role in food, beverage and nutraceutical processing for years to come.
4.8 Sources of further information and advice ∑
Ganetsos, G., Barker, P.E. Preparative and production scale chromatography. Marcel Dekker, Inc., 1992. ∑ Guiochon, G., Shirazi, S.G., Katti, A.M. Fundamentals of preparative and nonlinear chromatography. Academic Press, Inc., 2nd Edition, 1994. ∑ Sofer, G., Hagel, L. Handbook of process chromatography – a guide to optimization, scale-up and validation; Academic Press, London, 3rd Edition, 2001.
4.9 List of abbreviations ACE CEC CMCRS
angiotensin converting enzyme cation-exchange chromatography continuous-membrane chromatography reactor © Woodhead Publishing Limited, 2010
138 Separation, extraction and concentration processes CSS DHA DVB EBA EPA FB FDA GMP GOS HIPE HMOS IEC IgG LEC LF LP LTP MC MHRA MMM MT PAT PDE QA QC RCA RP-HPLC SEC SFC SMB SOP
combined stabilization system docosahexaenoic acid divinylbenzene expanded bed adsorption eicosapentaenoic acid fixed bed Food and Drug Administration good manufacturing practice galacto oligosaccharides high internal phase emulsion human milk oligosaccharides ion-exchange chromatography immunoglobulin G ligand-exchange chromatography lactoferrin lactoperoxidase lipid transfer protein multi-component Medicines and Healthcare products Regulatory Agency mixed matrix membranes mass transfer process analytical techniques partial differential equation quality assurance quality control ricinus communis agglutin reversed-phase high-performance liquid chromatography size-exclusion chromatography supercritical-fluid chromatography simulated moving bed standard operating procedure
4.10 References Abel, S., Babler, M.U., Arpagaus, C., Mazzotti, M., Stadler, J., (2004) Two-fraction and three-fraction continuous simulated moving bed separation of nucleosides, J Chromatogr A, 1043(2): 201. Adachi, S., (1994) Simulated moving bed chromatography for continuous separation of two components and its application to bioreactors, J Chromatogr, 658: 271. Alkio, M., Gonzales, C., Jäntti, M., Aaltonen, O., (2000) Purification of polyunsaturated fatty acid esters from tuna oil with supercritical fluid chromatography, J Am Oil Chem Soc, 77: 315–321. Al-Mashikhi, S.A., Li-Chan, E., Nakai, S., (1988) Separation of immunoglobulins and lactoferrin from cheese whey by chelating chromatography, J Dairy Sci, 71: 1747–1755.
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Advances in process chromatography and applications 145 Noh, K.M., Imm, J.Y., (2005) One-step separation of lysozyme by reverse micelles formed by the cationic surfactant, cetyldimethylammonium bromide, Food Chem., 93: 95–101. Noppe, W., Haezebrouck, P., Hanssens, I., De Ouyper, M., (1999) A simplified purification procedure of alpha-lactalbumin from milk using Ca 2+-dependent adsorption in hydrophobic expanded bed chromatography, Bioseparation, 8(1/5), 153–158. Ottens, M., Houwing, J., van Hateren, S.H., van Baalen, T. van der Wielen, L.A.M., (2006a) Multi-component fractionation in SMB chromatography for the purification of active fractions from protein hydrolysates, Food Bioprod. Process., 84(C1), 59. Ottens, M., Leene, G., Beenackers, A.A.C.M., Cameron, N., Sherrington, D.C., (2000) PolyHipe: a new polymeric support for heterogeneous catalytic reactions: kinetics of hydration of cyclohexene in two- and three-phase systems over a strongly acidic sulfonated PolyHipe, Ind. Eng. Chem. Res., 39: 259–266. Ottens, M., Wesselingh, J.A., van der Wielen, L.A.M., (2006b) Downstream processing in biotechnology, Chapter 9 in Basic biotechnology, edited by Ratledge and Kristiansen, 3rd edition, Cambridge University Press. Owen, R.O., Chase, H.A., (1997) Direct purification of lysozyme using continuous counter-current expanded bed adsorption, J. Chromatogr. A, 757: 41–49. Paananen, H., (1996) Trends in the chromatographic separation of molasses, SPRI 1996 Workshop on Separation Processes in the Sugar Industry, New Orleans, USA. Paananen, H., Kuisma, J., (1999) Multiple value-added products using the FAST separation technology, International Conference on Value-Added Products for the Sugar Industry, Baton Rouge, USA. Paillat, M.C.P., (1999) Different industrial applications of continuous chromatography in the sugar industry and for the production of sugar derivatives, Detmold Starch Convention, Detmold, Germany. Pais, L.S., Loureiro, J.M., Rodrigues, A.E., (1997) Separation of 1,10-bi-2-naphthol enantiomers by continuous chromatography in simulated moving bed, CES, 52(2): 245. Paredes, G., Makart, S., Stadler, J., Mazzotti, M., (2005) Simulated moving bed operation for size exclusion plasmid purification, CET, 28(11): 1335. Pearce, R.S., Houlston, C.E., Atherton, K.M., Rixon, J.E., Harrison, P., Hugues, M.A., Dunn, M.A., (1998) Plant Physiol., 117: 787. Pflegerl, K., Podgornik, A., Berger, E., Jungbauer, A., (2002b) Screening for peptide ligands on CIM monoliths, Biotechnol. Bioeng., 79: 733–740. Pflegerl, K., Podgornik, A., Schallaun, E., Jungbauer, A., (2002c) Direct synthesis of peptides on CIM monolithic columns for affinity chromatography, J. Combinat. Chem., 4: 33–37. Pflegerl, K., Hahn, R., Berger, E., Jungbauer, A., (2002a) Mutational analysis of a blood coagulation factor VIII-binding peptide, J. Peptide Res., 59: 174–182. Plate, K., Beutel, S., Buchholz, H., Demmer, W., Fischerfruhholz, S., Reif, O., Ulber, R., Scheper, T. (2006) Isolation of lactoferrin, lactoperoxidase and enzymatically prepared lactoferrin from proteolytic digestion of bovine lactoferrin using membrane adsorptive chromatography, J. Chromatogr. A, 1117, 81–86. Podgornik, A, Barut, M., Strancar, A., Josic, D., Koloini, T., (2000) Construction of large-volume monolithic columns, Anal. Chem., 15: 5693. Raab, B., Leman, H., Schwenke, K.D., Kozlowska, H., (1992) Comparative study of the protein patterns of some rapeseed (Brassica napus L.) varieties by means of polyacrylamide gel electrophoresis and high-performance liquid chromatography, Nahrung, 36: 239–247. Raab, B., Schwenke, K.D., (1984) Simplified isolation procedure for the 12 S globulin and the albumin fraction from rapeseed (Brassica napus L.), Nahrung, 8: 863–866. Rehmanji, M., Gopal, C., Mola, A., (2000) Tech. Q. Master Brew. Assoc. Am., 39: 24–28, A novel stabilization of beer with Polyclar Brewbrite.
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146 Separation, extraction and concentration processes Rito-Palomares, M., Miguel H., (1998) Influence of system and process parameters on partitioning of cheese whey proteins in aqueous two-phase systems, J. Chromatogr. B, 711, 81. Rodrigues, A.E., Pais, L.S., (2004) Design of SMB chiral separations using the concept of separation volume, Sep. Purif. Technol., 39(2): 245. Rousset. F., (1997) New developments in chromatographic separation of beet molasses. British Sugar EuroTechLink 97. Roy, I., Rao, M.V.S., Gupta, M.N., (2003) Purification of lysozyme from other hen’segg-white proteins using metal-affinity precipitation, Biotechnol. Appl. Biochem., 37: 9–14. Ruckenstein, E., Zeng, X.F., (1997) Macroporous chitin affinity membranes for lysozyme separation, Biotechnol. Bioeng., 56: 610–617. Safarik, I., Safarikova, M., (2003) Batch isolation of hen egg white lysozyme with magnetic chitin, J. Biochem. Biophys. Methods, 27: 327–330. Safarik, I., Safarikova, M., (2004) Magnetic techniques for the isolation and purification of proteins and peptides, Biomagn. Res. Technol., 2: 7. Salminen, S., Bouley, C., Boutron-Ruaultm M.C., Cummings, J.H., Franck, A., Gibson, G.R., Isolauri, E., Moreau, M.G., Roberfroid, M., Rowland, I.R. (1998) Functional food science and gastrointestinal physiology and function. Br. J. Nutr., 80(suppl): S147–71. Saufi, S.M., Fee, C.J., (2009) Fractionation of b-lactoglobulin from whey by mixed matrix membrane ion exchange chromatography, Biotechnol. Bioeng., 103: 138–147. Shinohara, Y., Kim, F., Shimizu, M., Goto, M., Tosu, M., Hasegawa, Y., (1994) Kinetic measurement of the interaction between an oligosaccharide and lectins by a biosensor based on surface plasmon resonance, Eur. J. Biochem., 223: 189–194. Spalding, B.J., ((1991) Downstream processing: key to slashing production costs 100 fold, Bio/Technology, 9: 229. Stahl, B., Thurl, S., Zeng, J., Karas, M., Hillenkamp, F., Steup, M., Sawatzki, G., (1994) Oligosaccharides from human milk as revealed by matrix-assisted laser desorption/ ionization mass, Anal. Biochem., 223: 218. Steffenson, M., Westerlund, D., (1996) J. Chromatogr. A, 720: 127–136. Su, C.K., Chiang, B.H., (2006) Partitioning and purification of lysozyme from chicken egg white using aqueous two-phase system, Process Biochem., 41: 257–263. Tejeda-Mansir, A., Montesinos, R.M., Magana-Plaza, I., Guzman, R., (2003) Breakthrough performance of stacks of dye–cellulosic fabric in affinity chromatography of lysozyme, Bioprocess Biosyst. Eng., 25: 235–242. Tennikova, T., Svec, F., Belenkii, B.G., (1990) High-performance membrane chromatography. A novel method of protein separation, J. Liq. Chromatogr. 13: 63–70. Thanh, V.H., Shibasaki, K., (1976) Proteins of soybean seeds. A straightforward fractionation and their characterisation, J. Agric. Food Chem., 24: 1117–1121. Thurl, S., Henker, J., Taut, H., Tovar, K., Sawatzki, G., (1993) Variations of neutral oligosaccharides and lactose in human milk during the feeding, Z. Ernahrungswiss., 32: 262–269. Thurl, S., Offermanns, J., Müller-Werner, B., Sawatzki, G., (1991) Determination of neutral oligosaccharide fractions from human milk by gel permeation chromatography, J. Chromatogr., 568: 291. Tomita, M., Wakabayashi, H., Yamauchi, K., Teraguchi, S., Hayasawa, H., (2002) Bovine lactoferrin and lactoferricin derived from milk: production and applications, Biochem. Cell Biol., 80(1): 109–112. Torre, M., Cohen, M.E., (1996) Perfusion liquid chromatography of whey proteins, J. Chromatogr. A, 729: 99–111. Van Walsum, H.J., Thompson, M.C., (1997) Simulated moving bed in the production of lysine, J. Biotechnol., 59: 127. Visser, S., Slangen, C.J., Rollema, H.S., (1991) Phenotyping of bovine milk proteins by
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Advances in process chromatography and applications 147 reversed-phase high-performance liquid chromatography, J. Chromatogr., 548(1–2): 361–370. Wang, X., Ching, C.B., (2005) Chiral separation of b-blocker drug (nadolol) by five-zone simulated moving bed chromatography, CES, 60: 1337–1347. Wei, C.H., Koh, C.J., (1978) Crystallographic characterization of a principal non-toxic lectin from seeds of Ricinus communis, Mol. Biol., 123, 707–711. Wilchek, M., Chaiken, I., (2000) An overview of affinity chromatography, Methods Mol. Biol. 147, 1–6. Wooley, R., Ma, Z., Wang, N.-H.L., (1998) A nine-zone simulating moving bed for the recovery of glucose and xylose from biomass hydrolyzate, IECR, 37: 3699–3709. Wu, D.J., Xie, Y., Wang, N.H.L., (1998) Design of simulated moving bed chromatography for amino acid separations, IECR, 37: 4023. Xu, Y., Sleigh, R., Hourigan, J., Johnson, R., (2000) Separation of bovine immunoglobulin G and glycomacropeptide from dairy whey, Process Biochem., 36: 393–399. Yamada, H., Fukumura, T., Ito, Y., Imoto, T., (1985) Chitin-coated celite as an affinity adsorbent for high performance liquid chromatography of lysozyme, Anal. Biochem., 146: 71–74. Yamasaki, N., Eto, T., (1981) A novel adsorbent for affinity chromatography of lysozyme, Agric. Biol. Chem., 45: 2939–2941. Ye, X., Yoshida, S., Ng, T.B., (2000) Isolation of lactoperoxidase, lactoferrin, a-lactalbumin, b-lactoglobulin B and b-lactoglobulin A from bovine rennet whey using ion exchange chromatography, Int. J. Biochem. Cell Biol., 32: 1143. Yoshida, S., (1988) Isolation of some minor milk proteins, distributed in acid whey from approximately 100,000 to 250,000 daltons of particle size, J. Dairy Sci., 71(1): 1–9. Yoshikawa, K., Okamura, M., Inokuchi, M., Sakuragawa, A., (2007) Ion chromatographic determination of organic acids in food samples using a permanent coating graphite carbon column, Talanta, 72: 305–309. Zydney, A.L., (1998) Protein separations using membrane filtration: new opportunities for whey fractionation, Int. Dairy J., 8, 243–250.
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5 Novel adsorbents and approaches for nutraceutical separation B. W. Woonton, CSIRO Food and Nutritional Sciences, Australia and G. W. Smithers, Food Industry Consultant, Australia
Abstract: In this chapter, an overview is provided of novel chromatographic adsorbents that are perhaps on or just over the horizon but hold promise as the basis for future industrial technologies in the cost-effective extraction of nutraceuticals from agri-food streams. Such adsorbents include molecular imprinted polymers, organic monoliths, stimuli-responsive resins, mesoporous molecular sieves, peptide affinity ligands, and membrane adsorbers. These adsorbents have the potential to improve specificity, selectivity, simplicity, robustness, and productivity, and to reduce the environmental impact of nutraceutical manufacture. Key words: nutraceuticals, separation, molecular imprinted polymers, monoliths, stimuli-responsive resins, molecular sieves, peptide affinity ligands, membrane adsorbers.
5.1 Introduction Nutraceuticals have entered the mainstream of foods and beverages. By definition, these components derived from agricultural and food raw materials have a positive influence on human health through a specific prophylactic or therapeutic physiological effect against chronic disease when consumed (Hardy, 2000). Nutraceuticals have usually been delivered to the consumer in a medicinal vehicle such as a tablet, capsule or in powdered form. More recently, the functional foods market, driven by consumer demands for healthpromoting foods and drinks, has seen nutraceuticals enter the food industry by way of the active component(s) in ‘functional’ varieties of everyday food and beverage products like yoghurt, bread, margarine, juice, and even water. This market is becoming a lucrative outlet for new and novel nutraceuticals, and © Woodhead Publishing Limited, 2010
Novel adsorbents and approaches for nutraceutical separation 149 food companies are keen to capture a slice of this market. These companies are thus exploring new sources of nutraceuticals, and the means to isolate, characterize and substantiate the bioactivity of such components to support the development of successful functional foods and beverages. In an earlier chapter (Chapter 4), more traditional approaches (e.g. fixed and moving bed chromatography) to the extraction, separation and concentration of nutraceuticals have been addressed. Chromatography is a powerful technique for the cost-effective separation and isolation of nutraceuticals from agrifood streams, including food processing co-product streams. Developments in continuous chromatography have enabled the commercial separation and isolation of biologically active proteins from milk and whey. For example, continuous separation (CSEP) technology is being used to isolate valuable dairy-derived molecules, including nutraceuticals (De Silva et al., 2003). Central to the chromatographic process is the relevant adsorbent that is responsible for effecting the separation. In this chapter, we introduce novel chromatographic adsorbents and approaches for the isolation of nutraceuticals, those that are perhaps on/over the horizon but hold promise as future industrial technologies for the cost-effective isolation of nutraceuticals, and the potential to deliver on consumer demands for effective functional foods and beverages. This overview is not intended to be exhaustive or inclusive of all novel adsorbents and technologies, but rather to provide a summary of the most promising adsorbents and approaches, that with further research and development can be employed for large-scale nutraceutical manufacture. The chapter provides an account of developments in: (i) (ii) (iii) (iv) (v) (vi)
molecular imprinted polymers; organic monoliths; stimuli-responsive resins; mesoporous molecular sieves; peptide affinity ligands; and membrane adsorbers.
5.2 Molecular imprinted polymers and applications in the nutraceutical industry Molecular recognition is an important process in many biological systems, for example the specific recognition of an antigen by an antibody. Such recognition mechanisms have been utilized in the development of biological affinity-based separation media (e.g. mono- or poly-clonal antibodies and affinity peptide resins) (Tharakan et al., 1992). Molecular imprinted polymers (MIPs) differ from these biological affinity systems in that they are synthetic organic materials designed with cavities specific to a template molecule employed during synthesis. In effect, they are tailored synthetic affinity materials that can be used in a range of applications including the separation © Woodhead Publishing Limited, 2010
150 Separation, extraction and concentration processes and analysis of molecules, the sensing of molecules, as artificial enzymes, and as antibody mimetics. MIPs are worthy of discussion as they have the potential to reduce the complexity of nutraceutical separation processes, thus reducing processing costs and the environmental impact associated with such processes. The concept of molecular imprinting was first reported by Kikuchi et al. (1972) and Wulff & Sarhan (1972), and since these early reports, there has been a significant amount of research undertaken to improve MIPs using alternative polymer chemistries and synthesis methodologies. There are a large number of publications reporting the development and application of MIPs for a range of target molecules including small organic molecules such as pharmaceuticals, pesticides, amino acids, peptides, high molecular weight proteins, polypeptides, steroids, sugars and cells (Andersson et al., 1995; Bolisay et al., 2006; Bossi et al., 2007; Nicholls et al., 1995; Rachkov & Minoura, 2001; Sellergren, 1994; Shi et al., 1999; Steinke et al., 1995). 5.2.1 Synthesis The synthesis of MIPs involves the organization of a cross-linked polymer matrix around a template molecule, followed by removal of the template, leaving recognition sites complementary to the template or similar molecules (Fig. 5.1). There are three main steps involved in the synthesis of a MIP, and these include: (i) pre-arrangement of selected monomers (such as vinyl or acrylic) around the target molecule; (ii) polymerization of the monomers (e.g. radical polymerization) with the addition of a cross-linker; and (iii) extraction of the target molecule leaving binding sites that recognize the target molecule (Gupta & Kumar, 2008). Two different approaches may be employed to manufacture MIPs, these are covalent and non-covalent imprinting. Non-covalent imprinting is the most common and involves the interaction of the template molecule with the monomers via hydrogen bonding, ionic bonding, hydrophobic interactions and van der Waals forces (Mosbach, 1994). After polymerization and removal of the template, the functional groups of the polymer are positioned to rebind the target molecule using the same non-covalent interactions. The non-covalent approach can only be employed where template molecules have functional groups capable of strong non-covalent interactions with the monomers. However, this can be overcome by using the covalent imprinting method, which involves covalently attaching the template molecule to a monomer followed by polymerization, cleavage of the covalent bond to liberate the template, and removal of the template. The reactive groups are positioned within the polymer cavities to allow covalent bonds to be reformed with the target molecule and provide selectivity. As there is a relatively small number
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(i) Complexation Template
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Fig. 5.1 Schematic illustration of the method for preparing molecular imprinted polymers (Takeuchi & Haginaka, 1999). Reproduced with permission. Copyright Elsevier (1999).
Novel adsorbents and approaches for nutraceutical separation 151
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Functional monomer
152 Separation, extraction and concentration processes of suitable reversible covalent bonding reactions that can be employed, the covalent imprinting method has limited applications (Gupta & Kumar, 2008). MIPs are often manufactured using bulk polymerization methods to produce monolithic materials which can be used for in situ separations, or ground and sieved to produce particles suitable for column packing and separation applications (Haginaka, 2009). MIPs can also be manufactured using suspension, seed and dispersion/precipitation polymerization methods to produce micro- or nano-spheres (Haginaka, 2009; Kempe & Kempe, 2006; Ye & Mosbach, 2001). Using such suspension polymerization methods, MIPs can be generated in aqueous environments; however, hydrogen bonding between the template molecule and the functional monomer, and therefore MIP selectivity, can be reduced when using aqueous solvents (Haginaka, 2009). Suspension polymerization techniques in liquid perfluorocarbon or mineral oil do not reduce hydrogen bonding and electrostatic interactions and the resultant MIP microspheres have been shown to be selective for biologically active molecules (Perez-Moral & Mayes, 2006). MIPs are usually packed into columns of various dimensions to allow (a) sample preparation before analysis; (b) removal of valuable molecules; or (c) removal of undesirable molecules from various raw material streams. 5.2.2 Advantages and disadvantages The benefit of MIPs is their stability in various chemical and physical environments for an extended period of time without any loss in template recognition (Ciardelli et al., 2006; Kriz & Mosbach, 1995). In addition, they are physically robust, resistant to elevated temperatures and pressures, and inert towards acids, bases, metals ions and organic solvents. The reported limitations of MIPs include binding site heterogeneity and slow mass transfer, which subsequently leads to peak broadening and tailing. In addition, MIPs prepared by bulk polymerization often have irregular sized and shaped particles which can lead to poor chromatographic performance and large pressure drops (Turiel et al., 2005). These limitations may be overcome by: (i) optimizing the MIP loading solvent; (ii) optimizing the MIP washing regimen to remove contaminant molecules; and (iii) finer control over the polymerization reaction leading to more regularly shaped particles. 5.2.3 Applications The characteristics of MIPs make them ideal for use in separation applications including chromatography, capillary electrophoresis, solid phase extraction and membrane separations. They also have potential to be used as microreactors, in immunoassays, as antibody mimics and artificial enzymes. Owing to their © Woodhead Publishing Limited, 2010
Novel adsorbents and approaches for nutraceutical separation 153 selectivity and robustness, MIPs have been applied in the selective isolation, concentration and analysis of food-derived molecules such as theophylline, caffeine, catechin, cholesterol, glycyrrhetinic, theanine, a-tocopherol and flavonoids (Haginaka, 2009). For example, Puoci et al. (2007) developed a MIP and optimized conditions for the selective extraction and analysis of a-tocopherol from bay leaf (Fig. 5.2). This method recovered approximately 60% of the a-tocopherol with high purity as measured by HPLC. Many other MIPs have been reported for the separation and analysis of agricultural herbicides, pesticides and fungicides, toxins, antibiotics and other chemicals in food (Haginaka, 2009). The ability to produce MIPs for the selective isolation of proteins and peptides is improving (Bossi et al., 2007; Rachkov & Minoura, 2001; Shi et al., 1999), and being employed to selectively target bacteria through their surface expressed proteins (Xue et al., 2009). 5.2.4 Commercialization MIPs can be employed in analytical applications as solid-phase extraction media to selectively increase the concentration of analyte, decrease sample preparation time and increase method sensitivity. MIPs for the solid phase extraction and analysis of food contaminants such as antibacterials (chloramphenicol, nitroimidazoles, fluoroquinolones) and herbicides (triazines) have been manufactured by MIP Technologies AB (Lund, Sweden) and distributed by Sigma-Alridch (St Louis, USA). MIPs for solid phase extraction and analysis of vitamins in food such as riboflavin have recently become available (MIP Technologies AB, Sweden). There are few reports in the public domain verifying large-scale applications of MIPs in the food, pharmaceutical and chemical industries. However, given the reported benefits of MIPs including stability and specificity, with further research and development into large-scale manufacture of these adsorbents together with scaled applications, it is likely that they will find targeted use in the isolation of nutraceuticals.
5.3 Organic monoliths and applications in the nutraceutical industry In contrast to columns packed with small particles to facilitate separation (e.g. silica or agarose beads), monoliths are continuous solid phases which can be synthesized in situ in confined spaces such as columns, capillaries, and microfluidic and microchip devices. Monoliths are unique as they have an interconnected skeletal structure with flow-through paths that provide very high permeability to a mobile phase, high flow rates, and low back pressures during chromatography. These characteristics make monoliths promising for the large-scale cost-effective purification of nutraceuticals
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154 Separation, extraction and concentration processes 0.9 0.8
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Fig. 5.2 HPLC chromatograms of (a) crude bay leaf extract and (b) bay leaf extract after isolation and concentration of the a-tocopherol (a-TP) (Puoci et al., 2007). Reproduced with permission. Copyright Elsevier (2007). © Woodhead Publishing Limited, 2010
Novel adsorbents and approaches for nutraceutical separation 155 in the food industry. Monoliths can be synthesized using either inorganic materials such as silica or organic materials such as methacrylate polymers. As silica has poor stability and operational robustness during the caustic cleaning procedures employed in the food industry, this chapter focuses on the development and application of organic polymeric monoliths. 5.3.1 Synthesis Polymeric monoliths were first introduced in the late 1980s (Hjerten et al., 1989; Svec & Fréchet, 1992). Since then there have been several publications detailing developments in their manufacture and use in liquid chromatography, capillary electrochromatography and microscale separations (Nordborg & Hilder, 2009). Polymeric monoliths are synthesized by mixing an initiator, organic monomers (including a cross-linking monomer) such as styrene, methacrylate, acrylamide-based materials and pore forming solvents (porogens). The reagents are placed within a suitable mould, such as a column, capillary or channel of a microfluidic device and polymerized using radical polymerization initiated by either thermal treatment, ultraviolet, gamma or microwave radiation (Nordborg & Hilder, 2009). As polymerization proceeds, the polymer chains formed become insoluble in the polymerization mixture, phase separation takes place and the polymer chains precipitate out of solution, forming the monolith. The properties of the monolith synthesized, including the size and distribution of macro-, meso- and micropores are determined by the composition of the polymer mixture and the polymerization conditions. For example, the mode of initiation, type and amount of cross-linker, porogen, and the polymerization temperature. The large number of monomers that can be chosen to impart charged groups, other monomers, cross-linking agents and porogens results in a very large number of variations that can be examined and optimized to produce the desired monolith (Nordborg & Hilder, 2009). Organic monoliths can also be synthesized using sub-zero temperatures, where ice-crystals act as the pore-forming elements. After formation of the monolith, the ice crystals are allowed to thaw, the resulting water removed, leaving behind macropores between 10 and 100 mm in diameter. Such monoliths are known as cryogels (Arvidsson et al., 2002, 2003). Solid granules such as sodium sulfate can also be employed during the synthesis to improve bi-model pore formation and the hydrolytic characteristics of the resulting polymeric monoliths (Fig. 5.3) (Du et al., 2007). 5.3.2 Functionalization To improve the applicability of polymeric monoliths for separation applications, functionalization either during or after polymerization is required. In the first approach, the choice of monomers with charged functional groups results in material having charged groups throughout the monolith, including its surface.
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156 Separation, extraction and concentration processes (a)
Acc. V Spot Magn Det WD 20.0 3.0 2000¥ SE 9.3
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Fig. 5.3 SEM image of a glycidyl methacrylate-co-ethylene glycol dimethacrylate monolith prepared in the presence of the organic porogens dodecanol and cyclohexanol, (a) without and (b) with sodium sulfate granules (Du et al., 2007). Reproduced with permission. Copyright Elsevier (2007).
Examples of charged monomers that can be employed include methacrylic acid and ethylene glycol dimethacrylate. Optimization of monolith chemistry must take into account that the charged monomers participate in the polymerization process and will not all be available for separation interactions on the surface © Woodhead Publishing Limited, 2010
Novel adsorbents and approaches for nutraceutical separation 157 of the monolith. The approach of post-polymerization modification has been examined and reported to be advantageous as the number of variables is reduced and the monolith chemistry can be optimized followed by modification of the surface with the desired functionality. Monoliths can be modified to achieve separation based on ion-exchange, affinity recognition, reversedphase, and hydrophobic interaction. 5.3.3 Advantages and disadvantages Polymer monoliths have similar benefits to their polymeric particles. Their pH and pressure stability are two advantages that allow them to be employed in a range of environments where high flow rates may lead to large pressure drops across the bed. However, as monoliths are made from one piece of material with large interconnected pores (Fig. 5.3), diffusion mass transfer is minimized and convective mass transfer is maximized. As convective mass transfer is faster than diffusive transport, higher mobile phase flow rates can be employed and faster separation of molecules can be achieved. Breakthrough curves of lysozyme and immunoglobulin G using a polymethacrylate monolithic column (Connective Interaction Media, BIA Separations; Ljubljana, Slovenia) and a column packed with Source 30S particles (GE Healthcare Life Sciences Uppsala, Sweden) have been studied (Hahn et al., 2002). As expected, mass transfer in the particle packed column was very dependent on the linear velocity of the mobile phase, molecule size, and feed concentration, whereas mass transfer in the monolith was not significantly affected by velocity, molecule size or feed concentration (Hahn et al., 2002). A disadvantage of monoliths is their low surface area as a result of their high porosity. Thus, monoliths have been reported to be more suitable for the isolation of large biomolecules (e.g. proteins) rather than small molecules. Monoliths also have a tendency to be hydrophobic and thus adsorb proteins via both ion-exchange and hydrophobic interaction. This characteristic makes proteins difficult to elute from the monolith without solvent which can lead to protein denaturation (Nordborg & Hilder, 2009). Research to develop polymeric monoliths with less hydrophobic characteristics or modification of monolith surfaces with hydrophilic coatings is being undertaken to improve their effectiveness in the isolation of biological molecules. 5.3.4 Applications Monoliths have been employed to separate a range of food-derived biomolecules and nutraceuticals. These include enzymes such as pectin methylesterase, endo-polygalacturonase (Vovk et al., 2005; Vovk & Simonovska, 2007), lysozyme (Svec & Fréchet, 1995), and proteins such as trypsin inhibitor (Svec & Fréchet, 1995), bovine serum albumin (Du et al., 2007), lactoferrin (Adam et al., 2008), casein peptides (Gu et al., 2006), and virus particles
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158 Separation, extraction and concentration processes (Kramberger et al., 2004). Hilder et al. (2004) has also demonstrated that monoliths coated with quaternary amine-functionalized latex particles can effectively separate sugars (Fig. 5.4), and separate glucose and maltose from amylase-hydrolyzed corn starch. There are a number of studies indicating that monolithic adsorbents have potential for the preparative-scale separation and purification of biomolecules, such as immunoglobulins (Brne et al., 2007), a-lactalbumin, enzymes, plasmid DNA, viruses and cells (Jungbauer & Hahn, 2008). In addition, monolithic columns have been evaluated for their potential to isolate immunoglobulins using simulated moving bed chromatography (Pennings et al., 2008). 5.3.5 Commercialization Currently, a number of companies are manufacturing and distributing polymeric monolithic columns for analytical applications. These include BIA Separations (Villach, Austria), Dionex (Sunnyvale, USA), and BioRad (Hercules, USA). BIA Separations are one of only a few companies designing and supplying organic monoliths for both small- and large-scale purification. Their disk format monoliths are manufactured for small-scale separations using axial flow design characteristics, and their tube format monoliths are prepared for 3 1
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Fig. 5.4 Separation of carbohydrates within 10 min using an optimized anionexchange latex-coated polymeric monolithic capillary column. Peaks: (1) d(+) galactose, (2) d(+)glucose, (3) d(+)xylose, (4) d(+)mannose, (5) maltose, (6) d(–) fructose, (7) sucrose (Hilder et al., 2004). Reproduced with permission. Copyright Elsevier (2004).
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Novel adsorbents and approaches for nutraceutical separation 159 larger scale separation using radial flow design characteristics. The latter design is employed for large columns (e.g. up to 8 L) and is reported to allow for a significant increase in the linear velocity through the monolithic bed, without adversely affecting performance or pressure characteristics. Optimization of surface chemistry and adsorption capacity is a prerequisite for the application of organic monoliths in large-scale isolation of nutraceuticals, together with a determination of the economic benefit of using such adsorbents over traditional particle-packed columns. As manufacturing processes are optimized, and costs reduced, larger monolithic columns will deliver cost-efficient nutraceutical separation and isolation processes in the food industry.
5.4 Stimuli-responsive resins and applications in the nutraceutical industry Polymers that undergo a significant change in their structure and behavior (e.g. solubility) in response to an external physical, chemical or electrical stimulus are termed ‘stimuli-responsive’, ‘smart’ or ‘intelligent’ polymers. A number of polymers have this property and various stimuli, including temperature, pH and light can be applied to induce a change in polymer conformation and subsequent properties. Stimuli-responsive polymers with temperature responsiveness have been extensively studied and characterized. These polymers have been grafted onto chromatographic supports to produce stimuli-responsive resins. These novel resins have the potential to reduce solvent usage, processing costs and the environmental impact of nutraceutical isolation in the food industry. The best characterized temperature-responsive polymer is poly(Nisopropylacrylamide (poly-NIPAAm). This polymer differs from traditional polymers as its solubility decreases as the temperature of the poly-NIPAAm solution increases. When a critical temperature is reached (32 °C for polyNIPAAm), hydrogen bonding of water to the amide group is reduced and the poly-NIPAAm becomes unstable, contracts and enters a globular state (Fig. 5.5). The temperature at which this occurs is called the lower critical solution temperature (LCST) (Ayano & Kanazawa, 2006). The LCST of temperature-responsive polymers including poly-NIPAAm can be modified by incorporating hydrophilic or hydrophobic monomers into the polymer (Ayano & Kanazawa, 2006). Hydrophilic monomers generally increase the LCST of the polymer whereas incorporation of hydrophobic monomers generally decrease the LCST of the polymer. 5.4.1 Synthesis Temperature-responsive resins are synthesized by grafting temperatureresponsive polymers to the surface of solid chromatographic supports. Silica © Woodhead Publishing Limited, 2010
160 Separation, extraction and concentration processes
H 2O PNIPAAm
+ Stimulus – Temperature
Below LCST
Above LCST
Fig. 5.5 Coil to globule transition and subsequent solution turbidity change when poly-NIPAAm (PNIPAAm) is heated above or cooled below the lower critical solution temperature (LCST) (Ayano & Kanazawa, 2006). Reproduced with permission. Copyright Wiley-VCH Verlag GmbH & Co. KGaA.
has been the predominant support employed using radical co-polymerization of initiator immobilized silica or activated ester amine coupling to aminopropylmodified silica (Maharjan et al., 2008). Temperature-responsive co-polymers that have been grafted onto silica beads for chromatography include polyNIPAAm, poly-NIPAAm-co-butyl methacrylate, poly-NIPAAm-co-acrylic acid, and poly-NIPAAm-co-acrylic acid-co-N-tert-butylacrylamide (Maharjan et al., 2008). The use of silica matrices for the separation of nutraceuticals in the food industry is limited owing to their lack of operational robustness, including their instability at the high pH employed during cleaning processes (Maharjan et al., 2008). A temperature-responsive resin employing a porous crosslinked agarose support was recently reported. This resin was synthesized by functionalizing agarose beads with the initiator 4,4-azobis(4-cyanovaleric acid), followed by polymerization of N-isopropylacrylamide, tert-butylacrylamide, acrylic acid and the cross-linking agent, N,N-methylenebisacrylamide onto the surface using free radical polymerization (Maharjan et al., 2009). 5.4.2 Applications Ion-exchange chromatography There have been a number of studies examining the effect of pH and temperature-responsive resins on retention and separation of biomolecules by ion-exchange chromatography. Temperature-responsive poly-NIPAAmco-acrylic acid-co-N-tert-butylacrylamide silica beads with cation-exchange functionality have been synthesized and employed to separate bioactive peptides and amino acids. This temperature-responsive silica resin showed greater retention and separation of basic bioactive peptides at higher temperatures
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Novel adsorbents and approaches for nutraceutical separation 161 than at lower temperatures when using aqueous conditions (Kobayashi et al., 2002; 2003). Similarly, silica beads grafted with poly-NIPAAm-cobutylmethacrylate-co-dimethylaminopropyl acrylamide show better separation of small bioactive molecules (anti-inflammatory compounds) at 50 °C than at 10 °C when using an isocratic aqueous mobile phase (Ayano et al., 2006). Although changes in the retention and separation behavior of molecules have been demonstrated with temperature responsive silica resins, there have been a minimal number of reports showing the effective capture and release of biological molecules. Research recently published using porous cross-linked agarose resins has demonstrated the potential of temperature-responsive cation-exchange agarose-based resins to selectively capture and isolate charged cationic proteins such as lactoferrin from mixed protein systems (Maharjan et al., 2009). The temperature-responsive agarose resin had a 3-fold higher lactoferrin adsorption capacity at 50 °C than at 20 °C (Fig. 5.6), and was able to selectively adsorb lactoferrin from a mixture of lactoferrin, a-lactalbumin and b-lactoglobulin. In addition, dynamic studies indicated that approximately 50% of the adsorbed lactoferrin could be eluted from the agarose by simply reducing the temperature. The remaining lactoferrin was eluted using low concentrations of NaCl (Maharjan et al., 2009). Hydrophobic-interaction chromatography A significant amount of research has been undertaken to develop and understand temperature-responsive resins for hydrophobic-interaction chromatography. For example, temperature-responsive poly-NIPAAm-co-butyl methacrylate functionalized silica beads have been employed to separate a mixture of insulin chains A and B, and b-endorphin fragment 1–27 (Fig. 5.7). Temperature Adsorbed lactoferrin (mg mL–1)
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Fig. 5.6 Equilibrium adsorption isotherms of lactoferrin onto poly-NIPAAm-coacrylic acid-co-N-tert-butylacrylamide agarose resin at 20 °C (continuous line) and 50 °C (dashed line) (error bars represent the standard deviation from the mean) (Maharjan et al., 2009). Reproduced with permission. Copyright Elsevier (2009).
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162 Separation, extraction and concentration processes 1, 2, 3 2
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Fig. 5.7 Separation of a mixture of insulin chains A and B, and endorphin fragment 1–27 on a poly-NIPAAm-co-butyl methacrylate modified silica resin. Column temperature: (a) 0 °C, (b) 40 °C. Mobile phase: aqueous 0.17M NaCl. Flow-rate: 0.5 ml min–1. Detection: 215 nm (Kanazawa et al., 1997). Reproduced with permission. Copyright Elsevier (1997).
gradients were able to alter the hydrophobicity of the resin, resulting in temperature-modulated peptide elution from the column (Kanazawa et al., 1997). In addition, temperature-responsive poly-NIPAAm silica beads have been employed to separate steroids with an aqueous mobile phase (Kanazawa et al., 1996). The steroids were unresolved at temperatures below the LCST and well resolved above the LCST. The improved resolution at temperatures above the LCST was most probably the result of increased hydrophobic interaction of the steroids with the poly-NIPAAm stationary phase at the higher temperatures. Size-exclusion chromatography Temperature-responsive polymers have also been applied to synthesize temperature-responsive size-exclusion resins. For example, poly-NIPAAm was grafted onto porous glass beads, the poly-NIPAAm end functionalized with mercaptopropionic acid, and the resulting resin employed for gel permeation size-exclusion chromatography (Gewehr et al., 1992). The elution time of the dextrans was substantially altered between 25 °C and 32 °C, possibly owing to a change in the effective pore size of the resin via transition of the poly-NIPAAm chains from coils to globules on the surface of the pores of the glass beads. Porous polystyrene beads grafted with poly-NIPAAm have also shown temperature-responsive size-exclusion characteristics (Hosoya et al., 1994). In this example, an increase in temperature prolonged the elution time of higher molecular weight dextrans, possibly owing to collapse of the poly-NIPAAm causing the pore size to expand, and thus permitting the dextrans to penetrate the pores.
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Novel adsorbents and approaches for nutraceutical separation 163 Affinity separations Temperature-responsive polymers can also be conjugated to proteins and employed as affinity separation materials. Cycling below and above the LCST of the temperature-responsive polymer induces reversible precipitation/ solubilization of the bioconjugates in the aqueous solution, thereby allowing separation of the bioconjugate and the target molecule. Proteins such as a-chymotrypsin (Kim & Park, 1998) and a-glucosidase (Hoshino et al., 1998) were purified using this method. In addition, temperature-dependent isolation of lactate dehydrogenase from porcine muscle was achieved using a dye-affinity agarose modified by physically adsorbing the smart polymer poly-N-vinylcaprolactam onto the surface. Using this method, the lactate dehydrogenase was adsorbed to the column at 40 °C and eluted by reducing the temperature to 23 °C and adding KCl (Galaev et al., 1994). 5.4.3 Commercialization Temperature-responsive polymers can be employed to separate a wide spectrum of molecules using a range of separation mechanisms (e.g. charge, size, affinity). However, temperature-responsive resins are not currently available commercially for large-scale separation applications. With further research and development into specific applications and scale-up, temperature-responsive resins with silica or other matrices could be employed commercially to isolate hydrophobic and charged nutraceutical molecules. Such processes should require less solvent (e.g. salt, ethanol), which would reduce the cost and environmental impact of such separations, and potentially improve retention of the biological activity of such isolated molecules.
5.5 Mesoporous molecular sieves and applications in the nutraceutical industry Mesoporous molecular sieves, also termed mobile crystalline materials (MCMs), are a family of inorganic solids with regular nanostructure, large internal surface area (>1000 m2 g–1) and void volume, and narrow pore size distribution. These remarkable materials, exhibiting defined and tailored pore diameters from approximately 2 nm through to 50 nm, offer new possibilities for designing highly effective adsorption media based upon their pore size and large internal surface area. Mesoporous molecular sieves, notably the MCM-4X family, were first reported in the early 1990s (Beck, 1991; Beck et al., 1992; Kresge et al., 1992), and have attracted considerable research attention since then as novel catalysts, guest supports in chemical reactions, and potential adsorbents for the isolation of biological components (Beck & Vartuli, 1996; Brady et al., 2008; Kisler et al., 2001a; 2001b; Selvam et al., 2001).
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164 Separation, extraction and concentration processes 5.5.1 Synthesis Mesoporous molecular sieves are prepared using a technique called liquidcrystal ‘templating’, illustrated schematically for MCM-41 in Fig. 5.8. In this process, mesoporous solids are synthesized by thermal treatment and ultimate calcination of aluminosilicate gels in the presence of suitable surfactants. Unlike other mesoporous inorganic materials, such as silica, the MCM family of mesoporous molecular sieves comprise regular arrays of non-intersecting uniform hexagonal channels (Fig. 5.8 and 5.9), the dimensions of which can be tailored within 1.5–50 nm through the selection of different surfactants (‘template molecules’) and reaction conditions (Gusev et al., 1996; Beck, 1991; Kresge et al., 1992). The pore size can be tailored during synthesis to accommodate biological molecules of almost any size (Fig. 5.9). Hexagonal array
Silicate
Calcination
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Fig. 5.8 Schematic illustration of the liquid-crystal ‘templating’ technique for manufacture of mesoporous molecular sieves such as MCM-41 (Kresge et al., 1992). Reproduced with permission of Macmillan Publishers Ltd: Nature. Copyright Macmillan Publishers Ltd (1992).
Fig. 5.9 Model of the molecular sieve MCM-41 illustrating methane and ethane inside one of the hexagonal pores (pore diameter ~3 nm).
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Novel adsorbents and approaches for nutraceutical separation 165 5.5.2 Advantages and disadvantages The high internal surface area, tailored pore size, and very narrow pore size distribution of these mesoporous molecular sieves are characteristics that make them ideal candidates as ‘designed media’ for size-exclusion and adsorptionbased separation of biological molecules, notably proteins, peptides and other functional ingredients (Brady et al., 2008; Kisler et al., 2001a,b). For these types of applications, the best studied mesoporous molecular sieves include MCM-41 and MCM-48 (Melo et al., 1999; Zhao et al., 1996). Although the uniform mesoporous pore size makes these materials attractive candidates for adsorption and separation of biological macromolecules, their stability in aqueous systems is limited (Kisler et al., 2001b). Such instability has been a major limitation in the widespread use of these materials for waterbased extractions, and in the separation and isolation of targeted bioactive components for use as nutraceuticals. In order to overcome this shortcoming, many researchers have modified the mesoporous materials with a hydrophobic coating to assist in stabilizing the material in water (Brady et al., 2008; Kisler et al., 2001b). Such modification of MCM-41, for example, has enhanced the utility of this mesoporous molecular sieve in biochemical applications (Brady et al., 2008; Kisler et al., 2001a; 2001b). These materials have also been studied as processing aids in water treatment and in the generation of potable water (Cooper & Burch, 1999). 5.5.3 Applications Size-based separations The utility of mesoporous molecular sieves as size-exclusion materials in the isolation and separation of proteins was first examined by Diaz and Balkus (1996) and subsequently by Washmon-Kriel et al. (2000) in studies of cytochrome C permeation and adsorption. This research showed a correlation between protein adsorption and the pore size of the molecular sieve being examined, as would be expected, and retention of biological activity, an important consideration when designing a separation protocol for isolation of bioactive components like nutraceuticals. Takahashi et al. (2000; 2001) and Vinu et al. (2004) also showed strong correlation between enzyme immobilization and pore size in a study of several macroporous molecular sieves including MCM-41. The isolation of both large and small biological molecules using MCM-41 and MCM-48 has been explored (Brady et al., 2008; Kisler et al., 2001a; 2001b; Xue et al., 2004). Trypsin, lysozyme and riboflavin have been studied as model biological solutes in order to examine the size-exclusion characteristics of these mesoporous materials and their potential as bioseparation media (Kisler et al., 2001a). Similarly, Brady et al. (2008) explored the potential of hierarchical mesoporous silica sieves in the isolation of valuable dairy bioactive components as functional food ingredients, and the feasibility of such materials as industrial-scale adsorbents.
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166 Separation, extraction and concentration processes Protein stabilization Stabilization of labile enzyme activity, both in aqueous and organic environments, may well be an important application for mesoporous molecular sieves. MCM-48 was used to immobilize penicillin G acrylase and to stabilize the activity of this enzyme (Xue et al., 2004), and was shown to be superior to MCM-41 for this application. Kim et al. (2006) provided a comprehensive review of the probable mechanisms involved in stabilization of enzymes by various nanostructures, including mesoporous molecular sieves, which include confinement, charge and hydrophobic interaction, and multipoint attachment to the protein. Wang & Caruso (2005) reported the usefulness of mesoporous silica and modified forms thereof as supports for immobilization and encapsulation of enzymes, and stabilization of biological activity, demonstrating success of the technology with catalase. Mesoporous sieves may also provide utility in expanding the application scope of enzymes for processing in organic solvents by stabilizing their catalytic activity in such environments (Takahashi et al., 2001; Yan et al., 2002). 5.5.4 Commercialization The highly regular structure, tailored and tuneable properties, uniform pore size distribution and high internal surface area of selected mesoporous sieves represent attractive properties for the design of modern separation materials and protocols based upon size-exclusion and surface-interaction chemistry. However, before their industrial potential can be realized and these characteristics exploited, the long-term stability and safety, effectiveness in processing agrifood streams, and scaleability and process economics must be established.
5.6 Peptide affinity ligands and phage display methodology and applications in the nutraceutical industry Affinity chromatography is a method for separating biochemical components, often macromolecules (proteins, peptides, nucleic acids), from complex mixtures based on a highly specific biological interaction such as that between an antigen and antibody, substrate and enzyme, or ligand and receptor. The rationale for this process was first reported in the early part of the 20th century (Starkenstein, 1910). However, it was not until the late 1960s, associated with developments in modern chromatographic resins and support materials, and chemistry/linking reactions for immobilization of functional ligands, that affinity chromatography emerged as a viable industrial technology for the isolation of valuable biological components, including nutraceuticals (Cuatrecasas & Anfinsen, 1971; Wilchek & Miron, 1999). Although affinity chromatography is a very powerful technique and has revolutionized the fields of modern biochemistry, biotechnology and medical science over the past 40 years, it is not without its challenges. These relate © Woodhead Publishing Limited, 2010
Novel adsorbents and approaches for nutraceutical separation 167 to the preparation of the affinity adsorbent (ligand and resin support) and the use of the adsorbent in processing, and include considerations of selectivity, functionality, stability, efficiency, capacity, repeatability, and productivity. The choice of functional ligand to be immobilized and to effect the affinity separation is vast, and is dictated by the characteristics of the molecule targeted for isolation. Such functional ligands include proteins, enzymes, peptides, nucleic acids, sugars, vitamins, lipids, synthetic organic molecules (e.g. dyes), and metal ions. Once the ligand is selected, immobilization chemistry, maintenance of biological functionality, stability, avoidance of steric effects associated with the resin support, and effectiveness of the adsorbent under processing conditions, among others, must be considered (Wilchek & Miron, 1999). In the design of modern affinity adsorbents, the selection of ligands with broad applicability that meet most of these requirements is important in addressing the cost-effectiveness consideration critical for industrial processes. In this regard, peptides offer an attractive option (Amatschek et al., 2000; Pflegerl et al., 2002). 5.6.1 Advantages and disadvantages The main advantage of small peptides is their high affinity and selectivity. In addition, they are usually stable to mild chemical, physical and biological conditions, and not prone to degradation by protease enzymes found in most agri-food streams. Peptides are also amenable to a variety of linkage chemistries that provide for a very stable covalent bond between peptide and resin support. Tailored peptides of almost any composition and sequence can be readily synthesized through modern solid-phase techniques. Parallel and combinatorial approaches to directed peptide synthesis can readily result in large peptide libraries which can be rapidly screened for their affinity to a target compound (Pflegerl et al., 2002). Although peptide affinity resins have high affinity and selectivity, they can have relatively low binding capacity, increasing the overall cost of nutraceutical manufacture and the cost of the final product. In addition, manufacturing specific peptides on a large scale can be time consuming and expensive. Research groups are currently working on methodologies to improve the binding capacity of peptide adsorbents and significant competition in the field of peptide production is rapidly improving manufacturing technologies and reducing costs. There are also concerns regarding the toxicity and biocompatibility of novel affinity peptides when employed to produce products for human consumption. Before peptide affinity ligands can be employed in the food industry, detailed studies to ensure that they are stable during processing and non-toxic will be required. 5.6.2 Peptide identification A technique developed recently that will perhaps revolutionize the development and selection of peptide ligands for affinity adsorbtion and other applications
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168 Separation, extraction and concentration processes is called phage display, based on a concept reported in the mid-1980s (Smith, 1985). Phage display is a powerful technique for the study of protein–protein, protein–peptide, and protein–nucleic acid interactions. This approach uses bacteriophage vectors to express a protein or peptide on the surface of the virus particle coded by genetic material residing on the inside of the particle. Phage display thus creates a connection between genotype and phenotype, and facilitates the creation of a library of variants of a peptide (or protein) expressed on the outside of the virus, and allows for rapid in vitro screening of these variants based upon their binding affinity to a target molecule (e.g. enzyme, antibody, receptor, cell structural protein). This in vitro selection process is called panning, and allows for the high-throughput screening of protein interactions (Kay et al., 2001; Koivunen et al., 1999). In its simplest form, panning is carried out by incubating a library of phage-expressed peptides with the target molecule, usually in a 96-well plate, washing away any unbound phage, and then eluting the specifically bound phage. The eluted phage is then taken through additional binding and amplification cycles to enrich the pool in favor of the specific binding peptide sequence (Kay et al., 2001). In this way, phage display/panning allows for the rapid discovery of useful peptide ligands with specific affinity for the target protein or other molecules (nutraceuticals), and thus the development of effective peptide affinity adsorbents for research and industrial applications. The phage display/panning technique has been widely used in drug discovery, notably for the identification of promising new ligands (e.g. enzyme inhibitors, and receptor agonists and antagonists) to target proteins. Perhaps the most successful commercial outcome from phage display/panning technology has been the therapeutic antibody drug adalimumab (Bain & Brazil, 2003; Scheinfeld, 2003). This antibody, developed by Cambridge Antibody Technology (Cambridge, UK), targets the inflammatory cytokine TNF-a and is marketed commercially by Abbott Laboratories as HUMIRA (‘human monoclonal antibody in rheumatoid arthritis’) (Scheinfeld, 2003). 5.6.3 Commercialization The use of phage display/panning in the discovery of specific peptide– biomolecule interactions that may lead to effective affinity adsorbents was simplified through the development and sale of commercial ‘kits’ containing pre-made random peptide libraries, as well as a cloning vector for construction of custom libraries. Such ‘kits’ and related reagents are available from a number of commercial biotechnology/analytical companies, including for example New England BioLabs (Ipswich, USA), V-Biolabs (Sawston, UK), Mimotopes (Melbourne, Australia), and Cambridge Antibody Technology (Cambridge, UK) (now part of AstraZeneca). Development and refinement of phage display technology by the pharmaceutical industry provides an example of a biotechnology that can be applied in food processing applications and nutraceutical isolation. Peptide
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Novel adsorbents and approaches for nutraceutical separation 169 affinity technologies are not yet commercially available for the isolation of nutraceuticals at commercial scale. However, the phage display technology represents a robust tool for research into novel peptide affinity adsorbents. With further research and development to optimize capacity and manufacturing procedures, peptide affinity resins aimed at cost-effective isolation and separation of nutraceuticals should become available in the near future.
5.7 Membrane adsorbers, membrane chromatography and applications in the nutraceutical industry A recent development in the arena of chromatographic adsorbents is that of membrane adsorbers aimed at improving the efficiency and simplicity of the chromatographic process by allowing the targeted separation to occur on and within a macroporous membrane support (see Charcosset, 1998; Klein, 2000; Przybycien et al., 2004 for reviews of alternative bioseparation processes including the use of membrane adsorbers). This process has also been termed membrane chromatography (Charcosset, 1998; Thömmes and Kula, 1995). Membrane adsorbers are porous membranes with functional ligands (such as ion-exchange or affinity groups) covalently attached to the membrane (Demmer et al., 1989) or immobilized within the membrane structure. 5.7.1 Advantages and disadvantages Membrane adsorbers have several potential advantages over traditional particle matrices and adsorbents (Charcosset, 1998). Ion-exchange and other functionalized membranes avoid problems with swelling or packing and can easily be scaled up. Membrane adsorbers have low compressibility characteristics and also tend to demonstrate lower back pressures than traditional chromatographic resin materials thus allowing for the rapid processing of large volumes of fluid. Large targeted molecules, such as proteins and large peptides, are convected through the membrane, rather than having to diffuse into a resin bead before adsorbing to the functional ligand, and thus diffusion limitations are minimized. Overall mass transfer is improved, leading to shorter cycle times and higher overall throughput (Brandt et al., 1988). These properties facilitate the use of processing, maintenance and cleaning protocols more diverse and adaptable to existing procedures in the food processing industry. Based on the advantages of membrane adsorbers over traditional chromatographic resins (Charcosset, 1998), these adsorbents, once scaled to industrial size, would appear to be best suited for the separation and isolation of targeted components, notably proteins and peptides, found in low concentrations in large volumes of liquid (Gebauer et al., 1997). Membrane adsorbers have similar disadvantages to traditional membrane technologies. For example, when processing protein
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170 Separation, extraction and concentration processes containing fluids, a secondary boundary layer can form, reducing flux rates and increasing membrane cleaning requirements. In addition, as membrane adsorbers are made from organic materials, they are sensitive to the elevated temperatures and cleaning regimes often used in the food industry. These membranes must therefore be cleaned using mild conditions and replaced when their adsorption capacity or flux rates fall outside critical process specification. As with traditional membranes, large pressure drops across the membrane should also be avoided as they can cause deformation and irreversible changes in pore-size characteristics. 5.7.2 Functionalization The functional ligand chemistry available with membrane adsorbers is almost endless and is similar to that for traditional chromatographic resins, including strong and weak anion and cation exchange, affinity interaction, reversed phase, and hydrophobic interaction. Selection of the appropriate ligand chemistry, and membrane support material and configuration, will depend upon the specific targeted component of interest and separation strategy, source and ultimate end-use application of the isolate. 5.7.3 Configurations Improvements in membrane materials, equipment for conducting membrane chromatography, and the means by which the targeted fluid is presented to the membrane give the user choice in designing the best combination of membrane adsorber, device, and geometric configuration to achieve the most cost-effective separation of the targeted compound. Macroporous functionalized membranes are available as single or stacked flat sheets, hollow fibers, spiral-wound sheets, dead-end/single-use filters, and cassettes from several established commercial manufacturers (e.g. Pall, Sartorius) and some new entrants in the field (e.g. Mosaic Systems, Natrix Separations). These configurations allow for simple dead-end filtration applications, but also the more commercially relevant cross-flow processing for the treatment of large volumes of liquid in a short timeframe. The benefits of membrane adsorbers are most apparent in flow-through applications because capacity constraints in retention mode, particularly at high loading rates, make resin chromatography a more attractive option for capture steps. 5.7.4 Applications Ion-exchange separations Ion-exchange (anion and cation) functionalized membrane adsorbers have been the most extensively researched as cost-effective alternatives to traditional resin chromatography for the isolation of valuable minor protein and peptide constituents from agri-food streams, notably dairy whey (Adisaputro et al.,
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Novel adsorbents and approaches for nutraceutical separation 171 1996; Mitchell et al., 1994; Yang et al., 1999; Zietlow & Etzel, 1995). Attention has been devoted to the use of strongly acidic membrane adsorbers in the isolation of lactoferrin and lactoperoxidase from cheese whey (Chiu & Etzel, 1997; Mitchell et al., 1994), although some attention has also been paid to the growth factor constituents (Mitchell et al., 1994). Glycosylated and non-glycosylated forms of the glycomacropeptide have been isolated from whey using anion-exchange membrane chromatography (Kreub & Kulozik, 2009; Kreub et al., 2008). In the former study, the authors describe a direct-capture method for the isolation of glycosylated glycomacropeptide at pilot scale using a membrane adsorption process. Membrane ion-exchange chromatography has also been described for the isolation of bovine serum albumin (He & Ulbricht, 2008; Sarfert & Etzel, 1997), soybean trypsin inhibitor (He & Ulbricht, 2008; Josic & Strancar, 1999), patatin from potato (Alt et al., 2004), a-lactalbumin (Yang et al., 2002), a large protein, thyroglobulin (Yang et al., 2002), and in the large-scale production of antibodies (Zhou & Tressel, 2005; 2006; Zhou et al., 2006). Affinity separations Integration of membrane processing and affinity chromatography by way of affinity membrane adsorbers provides a number of advantages over traditional affinity chromatography with resin packed columns, notably in respect to time and recovery of bioactivity (Brandt et al., 1988; Klein, 2000; Zou et al., 2001). Such membrane adsorbents have been applied to the isolation of pepsin and chymosin using an immobilized pepstatin A membrane (Suen & Etzel, 1994), and to the separation of specific antibodies using membraneimmobilized bovine serum albumin, soybean trypsin inhibitor, Protein G, and immunoglobulin G (Kochan et al., 1996; Ostryanina et al., 2002). Affinity membranes can be dominated by slow sorption kinetics thus potentially impacting the performance of the separation (Suen & Etzel, 1994). Many of the protein and peptide examples shown have the potential to serve as valuable nutraceuticals in functional food applications (Etzel, 2004; McIntosh et al., 1998), and as natural antibacterial and bioactive agents and thus enter areas of both human and veterinary medicine, and biotechnology (Dionysius et al., 1993; Perraudin, 1991; Pouliot & Gauthier, 2006; Regester & Belford, 1999; Smithers, 2004; 2008). Micro-organism removal Membrane adsorbers provide an attractive alternative to traditional resinbased chromatography columns used to remove trace impurities and microbial contaminants (e.g. virus particles) in downstream processing applications, notably those requiring good manufacturing practice (GMP) compliance (Etzel & Riordan, 2006; Zhou et al., 2008). These membranes have been reported to reproducibly achieve greater than 4-log removal of mammalian viruses and greater than 3-log removal of endotoxin and contaminant DNA (Zhou et al., 2008).
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172 Separation, extraction and concentration processes 5.7.5 Commercialization Pall Filtration (Port Washington, USA) and Sartorius (Goettingen, Germany) are perhaps the best known manufacturers of membrane adsorbers for research and small industrial use. These manufacturers have targeted agrifood, biotechnology and biopharmaceutical processing applications for their membrane adsorber products, primarily the ion-exchange offerings. There are a number of citations in the scientific literature on the effective use of membrane ion exchangers in the isolation of biological molecules (He & Ulbricht, 2008; Mitchell et al., 1994; Przybycien et al., 2004; Yang et al., 1999). Several new manufacturers have introduced further innovations into the field that improve the inherent efficiency and simplicity of membrane adsorbers. For example, Mosaic Systems (Ba Breda, The Netherlands) is a biotechnology enterprise with a specialization in commercial-scale separation and isolation of components for the food ingredient and biopharmaceutical industries. This company has developed a novel mixed-matrix membrane adsorber that comprises functionalized resin beads (5–10 mm) immobilized in a macroporous membrane. The company claims that the embedded resin beads retain full chromatographic functionality resulting in the products’ selectivity and binding capacity, but that the macroporous membrane decouples pressure drop from the resin particle size leading to enhanced hydrodynamic performance. This mixed-matrix system leads to improved throughput and productivity compared with an equivalent traditional chromatographic system, minimal fouling, reduced water and buffer use, and retention of efficiency and specificity through multiple adsorption cycles. Mosaic Systems claims their mixed-matrix adsorbers have been proven in trials with customers for the manufacture of food ingredients, biopharmaceuticals, and the isolation of valuable components from dairy streams. Natrix Separations (Burlington, Canada) (formerly Nysa Technologies) has developed a patented polymeric hydrogel technology that combines the high binding capacity, selectivity and specificity associated with traditional chromatographic resins with the high throughput and ease of use of macroporous membranes (Childs et al., 2008). The membrane adsorber consists of a polymeric hydrogel formed within a flexible porous support matrix. The matrix provides mechanical strength, whereas the hydrogel characteristics determine the separation chemistry of the product. The hydrogel polymer provides high binding site density, a large surface area for binding, and rapid mass transfer supporting high flow rates while providing highly efficient capture of the target molecule (Childs et al., 2008).
5.8 Conclusions and sources of further information and advice There appear to be many advantages to the application of ‘next-generation’ separation materials, such as those covered in this chapter, in the isolation © Woodhead Publishing Limited, 2010
Novel adsorbents and approaches for nutraceutical separation 173 of nutraceuticals. These advantages include improved specificity, selectivity, simplicity, robustness, productivity and reduced environmental impact. However, several of these materials do remain on or over the horizon from an industrial-scale perspective. Further research is needed to establish: (i) large-scale manufacturing procedures for a number of these novel adsorbents; (ii) specific scaled applications in nutraceutical manufacture; (iii) their processing advantages over traditional chromatographic materials and devices; and (iv) perhaps of most importance from an industrial perspective, process economics and commercial feasibility. The reader is directed to the following websites for further information about the novel adsorbents and approaches covered in this chapter. ∑ ∑ ∑ ∑ ∑ ∑
Molecular imprinted polymers: www.MIPtechnologies.com www.molecular-imprinting.com http://www.sigmaaldrich.com/analytical-chromatography/samplepreparation/spe/supelmip.html Organic monoliths: www.bio-labs.com www.BIAtechnologies.com www.dionex.com Stimuli-responsive polymers: http://macromolecules.case.edu/research_stimuli-responsive.htm Mesoporous molecular sieves: www.mrs.org Peptide affinity ligands and phage display methodology: www.neb.com www.vbiolabs.com www.mimotopes.com www.cambridgeantibody.com Membrane adsorbers and membrane chromatography: www.pall.com www.sartorius.com www.mosaicsystems.nl www.natrixseparations.com
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180 Separation, extraction and concentration processes
6 Advances in the effective application of membrane technologies in the food industry M. Pinelo, G. Jonsson and A. S. Meyer, Technical University of Denmark, Denmark Abstract: This chapter focuses on the recent advances in the use of membrane technology for efficient separation and concentration of solutes in the dairy and fruit juice industry, as well as advances in the purification of bioactive compounds to be used as food additives. The importance of fouling reduction is emphasized because this is necessary for membrane processes to become economically feasible. Key words: membranes, dairy industry, fruit juice, fouling, bioactive compounds, membrane bioreactors.
6.1 Introduction Membrane separation technologies have attracted much attention in the food processing industries over recent decades. The main reason for the interest is that membranes allow efficient separation and concentration of solutes without changing phase, maintaining worthy chemical and physical properties of food components and systems with particular economical relevance. This chapter primarily focuses on the use of membrane technology in the dairy and fruit juice industries. The incorporation of new, advanced membrane filtration techniques generally occurred first in the dairy industry, and was only later incorporated into other food manufacturing processes. Fouling is still the main challenge to overcome before use of membrane technologies becomes economically feasible in industrial processes. Various techniques to reduce fouling, related to the dairy industry, are discussed in the first section of the chapter. In the fruit juice industry, membrane technology is mainly used for concentration and clarification purposes. The main challenge is to
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Advances in the effective application of membrane technologies 181 reduce the volume of juice maintaining the aroma compounds, which are otherwise lost when concentration methods such as evaporation are employed. The most recent advances in reverse osmosis and membrane distillation are reviewed, as these techniques are the most promising ones to maintain aroma compounds and juice quality. Wastewater treatment in food processing is also discussed in the chapter, as regulation regarding purified water quality and landfill management are becoming more and more restrictive. We pay special attention to wastewater treatment by using membrane bioreactors (MBR), particularly anaerobic MBR, on which research has focused on the recent years. Membrane separation is also a promising down-stream technique for purification of food ingredients and additives. We present the current challenges and limitations of applying membrane separation, coupled or not with enzyme technology, to purification and fractionation of oligosaccharides, as an example of an integrated process where membrane technology can act in synergy with other unit operations and processes.
6.2 Theoretical fundamentals of membrane separation Implementation of membrane technology in the dairy and fruit juice industry is determined by the economical feasibility of the process, which is directly associated with the separation efficiency and impact of fouling. Several theories and models have been developed in order to relate separation efficiency to the separation set-up, molecular features of solute, solvent and membrane, and operational conditions during filtration. One of these models (Jonsson and Boesen, 1975; Jonsson et al., 2008) is based on the transport of solvent and solute through the membrane pores and the friction of the solute particles and the pore wall. The model relates separation efficiency to several physical parameters: bˆ cm Ê t l Jv ˆ Ê [6.1] = 1 = b + 1 – ˜ exp Á – cp 1 – R K ÁË K¯ Ë e Di ˜¯ where cm and cp are the solute concentrations on the membrane surface and at the permeate, respectively. R is the rejection factor, b is the friction coefficient (dependent on solute and pore diameter), K is the distribution factor between pore fluid and bulk solution, t is a tortuosity factor, l is the real thickness of the fouling layer, e is the fractional pore area, Jv is the volumetric flow per unit membrane area and Di is the diffusion coefficient for a component of the solute. In certain conditions, e.g. high fluxes, concentration polarization can have a particular relevance on membrane performance. There are equations to quantify the influence of concentration polarization (e.g. Jonsson, 1985):
cm – cp ˆ Ê = exp Á J v d ˜ Ë Di ¯ cb – cp
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[6.2]
182 Separation, extraction and concentration processes and
Di =
kT 3 ph di
[6.3]
where c b represents the solute concentration in the bulk solution, d the polarization layer thickness, k the Boltzmann’s constant, and h the viscosity. These models and equations can be used as a tool to determine preliminarily, within a certain error interval, the efficiency of a particular process that could eventually be implemented at a large scale.
6.3 Membrane technology in the dairy industry Use of membrane technology in the dairy industry represents ~40% of the total applications of membranes in food processes (Daufin et al., 2001). Both polymeric and ceramic membranes are used for concentration purposes in the dairy industry. Among the polymeric ones, spiral cartridges of polysulfone or polyethersulfone are usually chosen, and tubular ceramic membranes of zirconium or aluminum oxide are also widely used. Ultrafiltration is the most widely used membrane separation process in the dairy industry (D’Souza et al., 2005). Ultrafiltration is used, for example, for concentration of milk before the cheese-making process or to reduce shipping volumes as well as for whey concentration, resulting in a protein-enriched stream which is being examined for novel applications, and which may have potential antiinflammatory and anti-cancer properties. 6.3.1 Fouling as a major concern in the dairy industry Fouling is the major concern for the dairy industry nowadays. In recent years, research efforts have aimed at the development of ultrafiltration systems with enhanced life time that can efficiently reduce the concentration polarization and fouling (Nigam et al., 2008; Rice et al., 2009a). Concentration polarization is provoked by an increasing solute concentration on the membrane surface owing to the transport of solute during convective solvent flux, which is rejected and accumulated on the membrane surface. The accumulation results in a back-diffusion of solute into the solution in the retentate side, hence creating a concentration gradient. As a consequence, an initial drop of permeate flux is observed, followed by a further progressive decrease caused by fouling. During microfiltration of skim milk for removal of bacterial spores, Guerra et al. (1997) observed a peak of permeate flux at 0.25 bar of transmembrane pressure, followed by a continuous decrease of permeate flux at increasing pressures (Fig. 6.1). After reaching 1 bar, the pressure was diminished, resulting in lower permeate fluxes than the ones obtained during
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Advances in the effective application of membrane technologies 183 the increasing pressure phase. This confirmed the influence of fouling on the permeate flux. Although theoretically the flux stabilizes when the steady-state is reached, in practice, the flux continues to decrease slowly with continued filtration. Also, fouling reduces dramatically the useful life of membranes (Rice et al., 2009b). The cleaning treatment is a key factor influencing the whole plant productivity. 6.3.2 Cleaning of membranes in the dairy industry D’Souza et al. (2005) made a complete study about the causes of membrane fouling as well as the main cleaning processes used in the dairy industry. They classified the cleaning agents into six different categories: (1) alkalis, which hydrolyze proteins and carbohydrates; (2) acids, for dissolving salts and oxides; (3) enzymes, particularly proteases and lipases; (4) surfactants, which decrease surface tension and charge; (5) sequestrants, for removal of minerals; and (6) disinfectants, since one of the main targets of the cleaning process is also the reduction of microbial load. A cleaning process generally involves three steps: product removal from the membrane and water rinsing; use of a single or a combination of appropriate cleaning agents followed by water rising; and disinfection (Trägårdh, 1989). In most cases, an acid–alkali treatment is used, the concentration of each cleaning agent being a critical parameter for each cleaning situation (Tohammadi et al., 275 250 225 Flux (L h–1 m–2)
200 175 150 125 100 75 50 25 0 0.0
0.1
0.2
0.3
0.4 0.5 0.6 0.7 0.8 Transmembrane pressure (bar)
0.9
1.0
1.1
Fig. 6.1 Permeate flux versus transmembrane pressure during microfiltration of skim milk for removal of bacterial spores using a ceramic membrane.The arrows show the direction in which the transmembrane pressure was changed. Adapted from Guerra et al. (1997).
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184 Separation, extraction and concentration processes 2002). Moreover, temperature and mechanical aspects i.e. flow rate of the cleaning solution and applied pressure during the cleaning process have to be carefully selected for each particular situation (Bartlett et al., 1995; Bird et al., 1995). The election of a proper cleaning method is also an important factor to consider in nanofiltration and reverse osmosis processes, which also have a major presence in the dairy industry. Nanofiltration is currently used, for example, to remove salt from whey as well as for the recovery of lactose from the permeate of the ultrafiltration step. Reverse osmosis has proved to be useful for whey concentration and for obtaining concentrated lactose by removing salt and minerals. 6.3.3 Recent advances for fouling reduction in the dairy industry In recent years, patents and publications on membrane technology applied to the dairy industry have focused, almost exclusively, on fouling reduction by (a) proposing new cleaning processes, (b) by optimizing the existing ones or (c) by coupling devices to the membrane set-up particularly designed to reduce the fouling, e.g. rotation/vibration and electric fields (see Section 6.7). Some have reported the advantage of using rotating or vibrating systems to enhance permeate flux and membrane selectivity (Frappart et al., 2006; Genkin et al., 2006; Jaffrin, 2008), providing promising results ready to be extrapolated to large-scale processes. Al-Akoum (2006) found a strong correlation between the permeate flux and the vibration speed in ultrafiltration of soy milk at 25 °C. The permeate flux increased ~10 times when using disks with 6-mm vanes at 11,000 rpm compared with smooth disks at 2500 rpm. Genkin et al. (2006) doubled the critical flux when using transverse vibration on submerged hollow fiber membranes, and with combined axial and transverse vibration, a five-fold increase of critical flux was observed. Frappart et al. (2006) increased the permeate flux from 130 to 230 L h–1 m–2 during recovery of lactose from diluted milk when using a rotating disk with vanes at 2000 rpm compared with a smooth disk at 1000 rpm. The same authors compared the performance of a vibrating system to the rotating disk, concluding that the rotating system is more efficient in reducing the fouling, owing to its higher shear rate. A recent patent (Makardij-Tossonian, 2009) proposes the use of an enzyme solution propelled by injecting gas to synergistically combine the effect of the enzyme solution on the breakdown of the particles causing fouling with the mechanical de-fouling effect promoted by the gas. The patent was especially developed for dairy food and beverage processing. Other inventions are more focused on the use of cleaning solutions than on the use of mechanical systems to prevent fouling. Flemming and Skou (2006) patented a system to reuse the cleaning solution by filtering it through the membrane, providing good results for ultrafiltration of dairy products. Omprakash et al. (2008) found that a 0.2% caustic solution provided the best performance for cleaning of fouled membranes used for concentration of whey protein. Measurement of fouling is another critical factor which
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Advances in the effective application of membrane technologies 185 has also been the subject of some patents. Mickols et al. (2008) conceived a system for on-line measurement of fouling in spiral wound membranes with particular potential for the dairy industry. It is based on the difference between acoustic impedances on the outer surface for a cylindrical wall of a permeate collection tube. Other systems applied other physical principles to measure the degree and nature of fouling, but the high costs associated with their implementation on a large scale is still hampering utilization at an industrial level. It can be concluded that the implementation and degree of profusion of membrane technology into the dairy industry over the coming years will be dependent on the economic impact that the several presented strategies for fouling reduction can have on the overall economy of the process. The most promising strategies are: (1) mechanical removal of fouling by using, for example, vibrating systems; (2) optimization of the conditions of use of specific chemical agents; and (3) use of enzymes able to breakdown the structure of the particular compounds causing the fouling, e.g. proteases for whey solutions (Nigam et al., 2008).
6.4 Membrane technology in the fruit juice industry Membrane technology is mainly used for clarification and for concentration of fruit juices, replacing the traditional evaporation (Table 6.1). There are several advantages of membrane processes over evaporation in fruit juices manufacture: (1) a higher quality of the final juice because of the reduction in the loss of aroma and other compounds with nutritional value; (2) lower energy consumption; and (3) versatility of the equipment which can be used to treat different products (Jiao et al., 2004). Use of traditional vacuum evaporation for concentration of fruit juices usually results in generation of off-flavours, the loss of aroma compounds and, in some cases, color degradation. Despite the drawbacks mentioned, vacuum evaporation is still the most widely used technique for concentration of juices in industry (GEA, 2009). This is probably ascribable to the fact that evaporation permits higher concentration efficiency than membrane treatments, which also in general, have, higher operational costs. 6.4.1 Reverse osmosis Amongst the membrane separation techniques used for juice concentration, reverse osmosis is the one of major interest for industry. The main advantage
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186 Separation, extraction and concentration processes Table 6.1 Some examples of utilization of membrane technology for concentration and clarification of fruit juices Application
Type of filtration
Conditions
Reference
Concentration of apple juice
Reverse osmosis
Spiral wound cellulose acetate and polyamide membranes 21.1–26.7 °C
Chua et al., 2007
Concentration of apple juice
Reverse osmosis and ultrafiltration
Plate and frame cellulose acetate membrane 20 °C
Sheu and Wiley, 1983
Clarification of cactus pear juice
Ultrafiltration
Hollow fiber polysulfone membrane 25 °C
Cassano et al., 2007a
Concentration of cactus pear juice
Osmotic distillation
Microporous polypropylene Cassano et al., hollow fiber 28 °C 2007a
Clarification of kiwifruit juice
Ultrafiltration
Hollow fiber polysulfone membrane 20–30 °C
Cassano et al., 2007b
Concentration of pineapple juice
Osmotic evaporation
Polytetrafluoroethylene membrane 25 °C
Hongvaleerat et al., 2008
Concentration of Reverse osmosis blackcurrant juice
Tubular polyethersulfone membrane 25 °C
Banvolgyi et al., 2009
Concentration of noni juice
Osmotic distillation
Hollow fiber polypropylene Valdes et al., 30 °C 2009
Concentration of Membrane blackcurrant juice distillation
Hollow fiber polypropylene Kozak et al., membrane 15 and 19 °C 2009
Recovery of the Vacuum main pear aroma membrane compound (ethyl distillation 2,4-decadienoate)
Hollow fiber module of polypropylene 25 °C
Diban et al., 2009
is that the process is carried out at low temperatures, which results in higher retention of juice components and reduction in the consumption of energy. The efficiency of the process is highly determined by the operational variables and type of membrane. Polyamide and cellulose acetate are the most common materials used for reverse osmosis of fruit juices. Polyamide membranes used for concentration of apple juice have been found to be more resistant, and have provided higher retention of flavors and higher flux than the cellulose acetate ones, and similar results were obtained for other fruit juices (Chua et al., 1987; Sheu et al., 1983; Palmieri et al., 1990). A different configuration also seems to play a main role in the performance of reverse osmosis. The plate and frame configuration have, in general, a higher retention capacity for flavor aromas than the spiral wound, owing to the lower membrane package density and area (Chou et al., 1991). Increases in transmembrane pressure and decreasing temperatures favored the retention of flavor compounds such as ethyl-2-methyl butanoate, hexanol and hexanal in concentration of apple juice (Alvarez et al., 1998; Chou et al., 1991). There are no results
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Advances in the effective application of membrane technologies 187 of advances or a significant research focus on reverse osmosis applied to concentration of fruit juices in recent years. This is probably ascribable to the limited capacity of reverse osmosis to achieve high concentration levels. Jiao et al. (2004) set 25–30 °Brix as the maximum achievable concentration of fruit juice when reverse osmosis is used, the low efficiency ascribable to the osmotic pressure, versus ~80 °Brix that can be reached by multistep vacuum evaporation. This suggests the use of reverse osmosis as a pre-step in the concentration process, which is coupled, for instance, to a subsequent evaporation. This is already being implemented in commercial plants (Gadea, 1987), being energy-saving and resulting in a higher production capacity. 6.4.2 Membrane distillation Membrane distillation can be defined as the transfer of a target solute between two solutions subjected to different temperatures and separated by a hydrophobic membrane. The temperature gradient creates a vapor pressure difference between the two interfaces, resulting in a water flux from the high-temperature to the low-temperature side. The volatile compounds evaporate and diffuse and/or convect across the membrane from the feed side, and are condensed in the other side (Khayet et al., 2002; Lawson and Lloyd, 1997). The maximum achievable concentration when using membrane distillation was reported to be 60–70 °Brix, which is very close to the one reachable by traditional evaporation (Jiao et al., 2004). The selectivity of the membrane for each of the components in the feed solution is determined by the vapor–liquid equilibrium. Therefore, the highest permeation rate will correspond to the component with the highest partial pressure (Mulder, 1996; Bandini and Sarti, 1999). The pore size of the membrane used for membrane distillation are usually between 0.2 and 1.0 mm and hydrophobic materials such as polyvinyldifluoride and polytetrafluoroethylene are the most used polymers (Mulder, 1996). The material of the membrane must be carefully selected. If the affinity between the material of the membrane and the properties of the feed solution is too high, an undesired wetting of the membrane can occur (Mengual et al., 2004; Mulder, 1996). As expected, the flux rate across the membrane will increase with higher porosities and thinner membranes. A higher temperature difference between juice and water results in a flux increase, however, temperature cannot be increased indefinitely, as some of the solute components of fruit juice are thermosensitive. As expected, permeate flux only slightly decreases with increasing juice feed concentration compared with reverse osmosis and, in general, flux is higher than for reverse osmosis at high concentration ratios (Calabro et al., 1994). Depending on the method by which the vapor is recovered from the membrane pores, the membrane distillation process can be performed by using several configurations (Fig. 6.2): (a) direct-contact membrane distillation (DCMD) uses an aqueous solution for vapor recovery; © Woodhead Publishing Limited, 2010
188 Separation, extraction and concentration processes Membrane
Aqueous solution
Flow Aqueous solution
Sweep gas Vapor
Vacuum or Air gap
(a) Vacuum Aqueous solution
Aqueous solution
Sweeping gas
(b)
(c) Boundary layers
Cfeed
Cfm
Cpermeate Cpm
Tfeed
Tfm Tpm Tpermeate
Membrane (d)
Fig. 6.2 Membrane distillation processes employed for concentration of juices: (a) cross-sectional view of a hydrophobic membrane in contact with an aqueous solution illustrating the vapor–liquid interfaces in MD; (b) configuration of SGMD; (c) configuration of VMD; (d) schematic illustration of temperature (T) and concentration (C) profiles in VMD and SGMD during the concentration process; membrane–feed boundary layer (fm), membrane–permeate boundary layer (pm).
(b) osmotic membrane distillation (OMD), an osmotic medium; (c) vacuum membrane distillation (VMD) couples to a vacuum system; (d) sweeping gas membrane distillation (SGMD) uses a sweeping gas; and
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Advances in the effective application of membrane technologies 189 (e) air-gap membrane distillation (AGMD), a stagnant air gap plus a cold plate. Of these, DCMD has probably been the most studied configuration. About 5 years ago, a very limited number of studies on application of membrane distillation for juice concentration or aroma recovery were available (Calabro et al., 1994; Khayet et al., 2002; Lagana et al., 2000). However, the number of studies on this topic has dramatically increased recently. Calabro et al. (1994) succeed in concentrating orange juice by membrane distillation with total retention of sugars and organic acids, but a considerable degradation of vitamin C caused by oxidation and high temperature. Temperature seems to be the most important factor for process improvement in membrane distillation. Bui et al. (2004) reported a flux increase of 3–4 times when feed temperature was 15 degrees higher in glucose solutions of 30 to 60%. A reduction of the flux is normally caused by viscosity increase when the concentration is higher than 50%. In all the studies, however, membrane distillation is still coupled to a prior evaporation or membrane filtration step. Kozak et al. (2006) used OMD as the final step for concentration of must and apple juice in laboratory and in a large-scale installation. After reverse osmosis, grape and apple juice reached a concentration of 23 °Brix, which increased up to ~62 °Brix by using OMD at 20 °C. No advances in the use of new membrane materials or setups are being investigated for improving membrane distillation performance. Instead, most of the studies are focused on the optimization of the operating variables for each particular fruit, in an attempt to enhance the total soluble solids concentration in the particular juice. Temperature was the major factor in the concentration of ethyl 2,4decadienoate, a pear aroma compound, over feed flow rate, feed concentration and pressure when a commercial hollow-fiber polypropylene membrane was used for VMD (Diban et al., 2009). Changes in temperature were also reported to have a higher effect than feed flux in the concentration of volatile aroma compounds from blackcurrant juice (Bagger-Jørgensen et al., 2004). The low temperature (10 °C) favored the concentration of blackcurrant aroma esters compared with 45 °C by VMD. Kozak et al. (2009) reported an increase of 80% of the permeate flux by increasing the temperature by only four degrees (from 15 to 19 °C) between both sides of the membrane, confirming the dramatic effect of this variable in blackcurrant juice concentration. In this case, a microfiltration treatment was used for pre-concentration, and the membrane distillation contribute to concentrate the juice from 22 to 58.2 °Brix. Many recent works have also measured antioxidant capacity and content of phenolic compounds, which are reported to exert beneficial health effects (Del Rio et al., 2010), in juices subjected to membrane distillation to prove the advantages of this technique over the traditional methods. The antioxidant capacity and the content of phenols remained constant after the membrane distillation treatment in cactus pear juice, noni juice and other red fruit juices (Cassano et al., 2007a; Valdés et al., 2009; Koroknai et al., 2008). The practically non-existent patents on the application of membrane © Woodhead Publishing Limited, 2010
190 Separation, extraction and concentration processes distillation for concentration of fruit juices suggest that interest in this topic is still confined to academic circles. However, the higher concentration degree attainable by using membrane distillation compared with reverse osmosis suggests a major use of this technique for the future. Integrated processes As stated above, reverse osmosis and membrane distillation processes are commonly coupled to pre-evaporation or, to a micro/ultrafiltration step at industrial level. Micro- or ultrafiltration has been proved to be useful not only for clarification purposes, but also to reduce the viscosity of the juice by removing solids and pectin, easing the posterior membrane processes by increasing the permeate flux (Jiao et al., 2004). Johnson et al. (1996) developed a method for concentrating orange juice, combining ultrafiltration for clarification with traditional evaporation, reaching a concentration higher than 80 °Brix. A more complex method has been proposed by Álvarez et al. (2002) for concentrating apple juice, including a membrane bioreactor for clarification, reverse osmosis for pre-concentration, pervaporation to recover aromas and traditional evaporation as a final step. Enzyme treatment and ultrafiltration has become the most usual clarification treatment. It has been used for concentration of blackcurrant juice coupled with reverse osmosis in an integrated process (Banvolgyi et al., 2009). In turn, membrane filtration techniques themselves are sometimes coupled to other techniques with the aim, in most instances of reducing fouling. Fernandes et al. (2007) patented a hybrid method that combines nanofiltration and electrodialysis for concentration and fouling removal of grape must, respectively.
6.5 Membrane technology for treatment of wastewater in the food industry Again, the dairy industry is the first food processing industry incorporating the new advances in membrane technology applied to wastewater treatment. It has been estimated that the dairy industry generates 0.2–11 l of waste water per liter of milk (Daufin et al., 2001). Dairy industry effluents have been reported as one of the highest volume industrial effluents (Wheatley, 1990). Carbohydrates, lipids and proteins are the most abundant organic compounds in the effluents of the dairy industry with a reported polluting charge of 0.2–2.5 g l–1 biological oxygen demand (Daufin et al., 2001; Perle et al., 1995). In general, biological processes are more efficient and costeffective than the physicochemical methods (Vidal et al., 2000). The high energy requirements associated with aerobic treatment plants have turned the research focus toward anaerobic treatments in recent years, with lower sludge production and no aeration requirements (Demirel et al., 2005). However, the membrane aerobic bioreactor is still quite common for treatment of food-processing wastewater in industry (Abdulgader et al., 2007). Membrane © Woodhead Publishing Limited, 2010
Advances in the effective application of membrane technologies 191 separation substitutes the clarification step used in traditional wastewater treatment, giving several advantages: (1) a significant reduction in the amount of sludge generated, (2) the capacity to treat a major volume of water using less space, and (3) a higher quality of the resulting water (Daufin et al., 2001). 6.5.1 Aerobic membrane bioreactor A membrane bioreactor couples the activated sludge process with membrane separation. An efficient pre-treatment, to avoid excessive fouling of the membrane must precede the biological treatment. In general, either microfiltration or ultrafiltration are used to separate the sludge from the treated water. The membrane can be either submerged or external. The most usual set-up is a submerged hollow fiber or plate membrane ultrafiltration membrane (Daufin et al., 2001; Melin et al., 2006). Once again, fouling is the major factor to assure high performance. Fouling depends, besides other general factors such as membrane type and hydrodynamic conditions, on the presence of compounds that must be produced by microbial metabolism or added to the sludge, e.g. polyelectrolytes (Melin et al., 2006). 6.5.2 Anaerobic membrane bioreactor The main advantage of anaerobic treatment compared with aerobic is the dramatic reduction in the amount of produced sludge, which is at least five times lower. In addition, the anaerobic treatment results in the production of gas that can be used to generate energy to be used elsewhere in the plant and allows a degree of degradation of the organic matter higher than 80% (ArrosAlileche et al., 2008). Some authors reported the need to operate at subcritical fluxes to assure an efficient control of fouling (Hughes and Field, 2006). The selection of the membrane is generally based on a complete retention of the biomass and very low solute rejection, critical for good performance of the membrane (Jefferson et al., 2000). Among the most usual membrane materials for membrane bioreactors, polypropylene and polyethylene are being industrially used (Kubota and Mitsubishi). Because fouling is a major factor in these systems, cleaning is a critical step in these operations and it can be done by chemical or mechanical treatments. Back-flushing has been presented as one of the most innovative and efficient mechanical techniques to remove fouling (see Section 6.7).
6.6 New applications of membrane technology for the food industry: concentration and fractionation of saccharides Production of nutri-functional compounds, highly claimed by consumers over the past decade for being able to promote healthy properties of food, has © Woodhead Publishing Limited, 2010
192 Separation, extraction and concentration processes promoted the development of new techniques for production of high-purity compounds, to be used as supplements or additives in food. Prebiotics are a good example. In recent years, many research efforts have targeted the production of oligosaccharide prebiotics for elaboration of products, e.g. beneficial-health supplements or human-like infant milk formulas (Bode, 2009). In most instances, prebiotics are produced by enzymatic break down of a polysaccharide of vegetal origin, which produces an oligosaccharide with bioactive properties (Beine et al., 2008). Membrane technology is the most feasible strategy for purification in industrial manufacture of enzymatically produced oligosaccharides (Table 6.2). However, the lack of systematic understanding of how properties of the membrane and the oligosaccharide affect membrane filtration has limited to a large extent the industrial development of processes for production of prebiotics (Pinelo et al., 2009). Some of the factors related to design of the most advantageous set-up are described in the next section and the influence of certain operational variables which have to be carefully considered to eventually exploit the process at an industrial level is explored. 6.6.1 Design of separation set-up In most small-scale studies, membrane filtration was performed in a batch dead-end system, which is the simplest set-up to operate (Pinelo et al., 2009). The system consists of a tank for the feed solution, which passes across the membrane module. The permeate is recovered in a proper reservoir (Swennen et al., 2005). A magnetic stirrer is commonly used to reduce fouling and to improve the reaction rates on a laboratory scale where the enzymatic reaction and the membrane filtration occur simultaneously (Sanz et al., 2005). Industrially, further research into other types of efficient mixing is required. In the laboratory, Sarney et al. (2000) used a membrane reactor equipped with a magnetic stirrer for the recovery of oligosaccharides from milk using a combined b-galactosidase treatment with nanofiltration. However, when fouling is a major factor for productivity, a cross-flow filtration system is recommended (Ramaswami et al., 2005). In cross-flow, the feed solution flows parallel to the membrane surface, sweeping the retained particles and helping to reduce the concentration polarization. This is probably the most advisable set-up for industrial purposes. A cross-flow system was used for purification of the products resulting from autohydrolysis of rice hulk in a tubular membrane (Vegas et al., 2006). Goulas et al. (2002) reported the same efficiency for dead-end and cross-flow filtration modes in the removal of glucose from enzymatically obtained oligosaccharides, but concluded that the concentration polarization was a major factor limiting the performance of the dead-end system.
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© Woodhead Publishing Limited, 2010
Substrate
Type of separation
Commercial powder of oligosaccharides from chicory rootstock
UF for separation of larger Oligosaccarides from UF cutoff: 20 kDa, impurities and NF for removal of 3 to 10 degrees of NF cutoff: 0.5–5 kDa monosaccharides polymerization (DP)
Kamada et al. (2002)
Arabinoxylan hydrolysates Commercial saccharides
Purification by UF (5, 10 and 30 kDa) Removal of di- and monosaccharides Purification of oligosaccarides and removal of lactose by NF Simultaneous concentration and purification by 1 kDa by NF Separation of di- and monosaccharides by NF
Stirred cell at 4 bar
Swennen et al. (2005)
Stirred cell at 50 °C and 4 bar
Sanz et al. (2005)
Caprine milk Liquors from rice husk autohydrolysis Commercial mixture of galacto-oligosaccharides
Product
Arabinoxylooligosaccharides Oligosaccharides Milk oligosaccharides Xylooligosaccharides Raffinose, sucrose and fructose
Commercial mixture of Concentration and purification by Oligosaccharides of saccharides from yacon UF and NF DP from 3 to 10 rootstock Hemicellulose hydrolyzate Purification by NF Xylose stream Enzymatic hydrolytes from vegetal origin Commercial oligosaccharide mixture Liquors from almond shells autohydrolysis
Concentration by NF Purification by NF Separation of lignin-related impurities by UF
Xylooligosaccharides Fructooligosaccharides Xylooligosaccharides
Filtration conditions
Reference
Feed velocity: 80–120 mL, 20–40 bar Sarney et al. (2000) Vegas et al. (2006) TiO2/ZrO2 ceramic membrane at 30 °C and 2–14 bar Feed concentration varied between Goulas et al. (2002) 0.15–0.08 g mL–1, filtration was performed at 25–60 °C and 5–30 bar UF cutoff 10 kDa (0.52 bar) and NF Olano-Martin et al. (2001) cutoff 5 kDa (5 bar) Concentrated feed solution: 20 wt% Retentate flow: 6 L min–1, filtration performed at 40–60 °C and 20–40 bar Feed velocity: 400 mL min–1 pH 5, 48 °C and 16 bar Feed flow: 1850 L h–1, filtration performed at 45 °C and 1 MPa Cutoffs between 1 and 8 kDa Filtration performed at 25 °C and 2.6–9 bar
Machado et al. (2000)
Huang et al. (2001) Espinoza-Gomez et al. (2004), Grassin et al. (1996) Li et al. (2004)
Advances in the effective application of membrane technologies 193
Table 6.2 Some examples of the use of ultrafiltration (UF) and nanofiltration (NF) membranes for the concentration and purification of oligosaccharides of different nature
194 Separation, extraction and concentration processes 6.6.2 Operational modes The integration of enzyme reactor and membrane is normally advantageous for recovery of prebiotic oligosaccharides from vegetal matrixes, mainly because this disposition diminishes product inhibition (Pinelo et al., 2009). The operational conditions for this particular integrated system depend to a large extent on whether the system operates with immobilized or free enzymes. On some occasions, the membrane itself was used for the immobilization. Nishizawa et al. (2000) reported the chemical immobilization of b-fructofuranosidase with glutaraldehyde to the inner surface of a ceramic membrane for production of fructo-oligosaccharides. Although the use of ceramic membranes has been used for purification of oligosaccharides by ultra- and nanofiltration, it is more common to use polymeric organic membranes for oligosaccharide purification (Artug et al., 2007). Cellulosic membranes seem more suitable to this purpose, being hydrophilic, but the use of polysulfone membranes is also common (Akhtar et al., 1995). A chemical surface treatment was recommended for a polysulfone membrane, in order to make it more hydrophilic (Akhtar et al., 1995). The operating variables that had a major influence on membrane filtration of oligosaccharides were: (1) temperature, (2) pH, (3) concentration of the feed solution, and (4) pressure. Temperature It has been reported that the permeate flux almost doubled when the temperature of saccharide solutions increased from 20 to 50 °C. This is ascribable to the reduction of viscosity, which favors the pass of the flow of the feed solution across the membrane (Goulas et al., 2002; Machado et al., 2000). The favorable effect of the temperature increase on the permeate flux was confirmed by Sjöman et al. (2008), during recovery of xylose from hemicellulose hydrolyzate feeds using various polysulfone membranes. pH For certain types of membrane carrying positive or negative charged groups, e.g. polycarbonate and sulfonated polysulfone membranes, the influence of pH on the saccharide rejection can be dramatic (Lazaridou et al., 2007). When the feed solution includes some charged compounds, flux commonly increases when working close to the isoelectric point of the membrane, as expected (Childress and Elimelech, 1996). Variation of pH can be used as a strategy for more efficient separation of charged saccharides, such as pectin, whose charge density and distribution is a function of the carboxyl groups in the structure.
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Advances in the effective application of membrane technologies 195 Concentration of the feed solution Permeate flux is expected to decrease with increasing feed concentration of the solution. As a result, a decrease in the rejection, at constant pressure, is usually observed in these cases (Goulas et al., 2002). Pressure The role of pressure and the recommended levels of pressure depend on the type and properties of the solution to be filtered. Generally, the rejection increases with increasing levels of pressure until reaching a certain level. This is ascribable to the higher flux of solute across the membrane, which decreases the difference in concentration between the solutions at both sides of the membrane. When pressure continues increasing, the concentration polarization and, in turn, the concentration of solute on the surface of the membrane increases too. As a consequence, the concentration of solute on the permeate side increases and the rejection decreases. In purification of xylo-oligosaccharides from lignin impurities, Nabarlatz et al. (2007) observed a decrease in rejection when the pressure was increased from 2.6 to 9 bar. It is then advised to operate at low pressures, within a range in which the concentration polarization is not so high, in order to maximize selectivity.
6.7 Future trends Implementation of membrane technology in food industry has been hampered by the economical limitations derived from fouling over recent decades. Many studies have been devoted to develop techniques and processes to reduce the impact of fouling and it seems that progress in membrane technology will be linked to progress in fouling reduction over the next few years. In particular, there are at least three techniques that have provided promising results for their efficiency and simplicity to be incorporated in a large-scale process: (1) high frequency backflushing, (2) vibrating membrane modules, and (3) electrofiltration. Some of the benefits associated with coupling these techniques in the dairy industry have been discussed in Section 6.3 but the applications in other food processes are increasing 6.7.1 High frequency backflushing Backflushing consists of applying pressure pulses from the permeate to the retentate side, and stripping off the particles causing fouling from the membrane surface into the bulk solution at the permeate side again (Pinelo et al., 2009). Backflushing is commonly achieved by placing a valve in the © Woodhead Publishing Limited, 2010
196 Separation, extraction and concentration processes 300
Flux (L h–1 m–2)
250 With backflush
200 150 100 50 0
Without backflush 0
15
30
45
60 75 Time (min)
90
105
120
Fig. 6.3 Influence of backflushing on the permeate flux versus time applied to the microfiltration of skim milk for removal of bacterial species using a ceramic membrane. Adapted from Guerra et al. (1997).
permeate side. The efficiency of backflushing depends both on the duration of the pulse itself and time between pulses. Guerra et al. (1997) reported a considerable increase of permeate flux when backflushing was incorporated to microfiltration of skim milk, as backflushing contributed to the removal of spores from the membrane surface (Fig. 6.3). Jonsson (2008) narrowed the intervals of fractionation of dextrans in a hollow-fiber ultrafiltration system using times between pulses from 1 to 30 s and backflushing times from 0.1 to 5 s. Application of backflushing in industry is still limited but further research in the coming years will permit a more detailed analysis of the balance between costs associated with implementation and economic benefit from fouling reduction. 6.7.2 Vibrating membrane module A vibrating design consists of applying high shear rates to favor the transmission of macromolecules through the membrane (Beier et al., 2007). This design proved to be particularly efficient for nanofiltration of diluted milk and for enzyme separation applied to food processes (Beier et al., 2007; Jaffrin et al., 2004). This system has been already presented in the section dedicated to fouling reduction in the dairy industry, but it is fair to include it here as a future trend technique, as many studies are currently being conducted to optimize the mechanical characteristics of the system for a feasible high performance of membrane separation in the dairy industry.
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Advances in the effective application of membrane technologies 197 6.7.3 Electrofiltration Electrofiltration is used when the particles causing concentration polarization have a net charge and can be attracted or repelled by applying an electric field, in this instance across the membrane (Pinelo et al., 2009). This technique has been successfully employed for separation of industrial solutions of enzymes with a surface charge, resulting in a dramatic increase of the permeate flux (Enevoldsen et al., 2007). The technique also gave good results in ultrafiltration of other proteins (Oussedik et al., 2000) and shows promise for purification of potential food ingredients in which charge can be induced, e.g. modifying the number and methoxyl groups on pectin molecules.
6.8 References Abdulgader M E, Yu Q J, Williams P, Zinatizadeli A A L (2007), ‘A review of the performance of aerobic bioreactors for treatment of food processing wastewater’. Proceedings of the International Conference on Environmental management, Engineering, Planning and Economics, pp. 1131–1136. Skiathos. Akhtar S, Hawes C, Dudley L, Reed P, Strafford P (1995), ‘Coatings reduce the fouling of microfiltration membranes’, J Membr Sci 107, 209–218. Al-Akoum O, Richfield D, Jaffrin M Y, Ding L H, Swart P (2006), ‘Recovery of trypsin inhibitor and soy milk concentration by dynamic filtration’, J Membr Sci 279, 291–300. Álvarez V, Riera F A, Álvarez S, Álvarez R (1998), ‘Permeation of apple aroma compounds in reverse osmosis’, Sep Purif Technol 14, 209–220. Álvarez S, Riera F A, Álvarez R, Coca J (2002), ‘Concentration of apple juice by reverse osmosis at laboratory and pilot-plant scale’, Ind Eng Chem Res 41, 6156–6164. Arros-Alileche S, Merin U, Daufin G, Gesan-Guiziou G (2008), ‘The membrane role in an anaerobic membrane bioreactor for purification of dairy wastewaters: A numerical simulation’. Bioresour Technol 99, 8237–8244. Artug G, Roosmasari K, Richau K, Hapke A (2007), ‘A comprehensive characterization of commercial nanofiltration membranes’, Sep Sci Technol 42, 2947–2986. Bagger-Jørgensen R, Meyer A S, Varming C, Jonsson G (2004), ‘Recovery of volatile aroma compounds from black currant juice by vacuum membrane distillation’, J Food Eng 64, 23–31. Bandini S, Sarti G C (1999), ‘Heat mass transport resistances in vacuum membrane distillation per drop’, AIChE J 45, 1422–1433. Banvolgyi S, Horvath S, Stefanovits-Banyai E, Bekassy-Molnar E, Vatai G (2009), ‘Integrated membrane process for blackcurrant juice concentration’, Desalination 241, 281–287. Barlett M, Bird M R, Howell J A (1995) ‘An experimental study for the development of a qualitative membrane cleaning model’, J Membr Sci, 218, 107–116. Beier S P, Jonsson G (2007), ‘Separation of enzymes and yeast cells with a vibrating hollow fiber membrane module’ Sep Purif Technol 53, 111–118. Beine R, Moraru R, Nimtz M, Na’amnieh S, Pawlowski A, Buchholz K, Seibel J (2008), ‘Synthesis of novel fructooligosaccharides by substrate and enzyme engineering’, J Biotechnol 138, 33–41. Bird M R, Barlett M (1995), ‘CIP optimization for the food industry: Relationships between detergent concentration, temperature and cleaning time’, Trans IChemE, 73, 63–70.
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198 Separation, extraction and concentration processes Bode L (2009), ‘Human milk oligosaccharides: prebiotics and beyond’, Nutr Rev 67, 183–191. Bui V A, Nguyen M H, Muller J (2004), ‘A laboratory study on glucose concentration by osmotic distillation in hollow fiber modules’, J Food Eng 63, 237–245. Calabro V, Jiao B L, Drioli E (1994), ‘Theoretical and experimental study on membrane distillation in the concentration of orange juice’, Ind Eng Chem Res 33, 1803– 1808. Cassano A, Conidi C, Timpone R, D’Avella M, Drioli E (2007a), ‘A membrane-based process for the clarification and the concentration of the cactus pear juice’, J Food Eng 80, 914–921. Cassano A, Donato L, Drioli E (2007b), ‘Ultrafiltration of kiwifruit juice: Operating parameters, juice quality and membrane fouling’, J Food Eng 79, 613–621. Childress A E, Elimelech M (1996), ‘Effect of solution chemistry on the surface charge of polymeric reverse osmosis and nanofiltration membranes’, J Membr Sci 119, 253–268. Chou F, Wiley R C, Schlimme D V (1991), ‘Reverse osmosis and flavor retention in apple juice concentration’, J Food Sci 56, 484–487. Chua H T, Rao M A, Acree T E, Cunningham D G (1987), ‘Reverse osmosis concentration of apple juice: flux and flavor retention by cellulose acetate and polyamide membranes’, J Food Proc Eng 9, 231–245. D’Souza N M, Mawson A J (2005), ‘Membrane cleaning in the dairy industry: a review’, Crit Rev Food Sci Nutr, 45, 125–134. Daufin G, Escudier J P, Carrere H, Berot S, Fillaudeau L, Decloux M (2001), ‘Recent and emerging applications of membrane processes in the food and dairy industry’, Trans IChemE, 79, 89–102. Del Rio D, Costa L G, Lean M E J, Crozier A (2010), ‘Polyphenols and health: what compounds are involved?’, Nutr Metab Cardiovasc Dis, 20, 1–6. Demirel B, Yenigun O, Onay T T (2005), ‘Anaerobic treatment of dairy wastewaters: a review’, Process Biochem 40, 2583–2595. Diban N, Voinea O C, Urtiaga A, Ortiz G (2009), ‘Vacuum membrane distillation of the main pear aroma juice compound. Experimental study and mass transfer modelling’, J Membr Sci 326, 64–75. Enevoldsen A D, Hansen E B, Jonsson G (2007), ‘Electro-ultrafiltration of industrial enzyme solutions’, J Membr Sci 28–37. Espinoza-Gomez H, Lin S W, Rogel-Hernandez E (2004), ‘Nanofiltration membrane pore diameter determination’, Rev Soc Quim Mex 48, 15–20. Fernandes G V M, Lopes C G, Correia N (2007) ‘Concentration and rectification of grape must involves combining nanofiltration and electrodialysis of grape must in hybrid process’, World Patent WO2008051100-A2. Flemming S, Skou F (2006), ‘Membrane filtration of a product, e. g. milk, in membrane plant, involves recovering cleaning solution after multi-step cleaning of membrane system in plant; and using recovered cleaning solution for cleaning membrane system’, European Patent EP1726353. Frappart M, Akoum O, Ding L H, Jaffrin M Y (2006), ‘Treatment of dairy process waters modeled by diluted milk using dynamic nanofiltration with a rotating disk module’, J Membr Sci 282, 465–472. Gadea A (1987), ‘Reverse osmosis of orange juice’. In Proceedings of the International fruit juice congress, Orlando, USA. GEA (2009) Report on ‘Evaporation technology for the juice industry’, GEA Wiegand GmbH, www.gea-wiegand.com. Accessed on Dec 15, 2009. Genkin G, Waite A G, Fane S, Chang S (2006), ‘The effect of vibration and coagulant addition on the filtration performance of submerged hollow fiber membranes’, J Membr Sci 281, 726–734. Goulas A K, Kapasakalidid P G, Sinclair H R, Rastall R A, Grandison A S (2002),
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Advances in the effective application of membrane technologies 199 ‘Purification of oligosaccharides by nanofiltration’, J Membr Technol 209, 321– 335. Grassin C, Fauquembergue P (1996), ‘Fruit Juices’, in: T. Godfrey, S. West (Eds.), Industrial Enzymology, Macmillan Press. Guerra A, Jonsson G, Rasmussen A, Nielsen E W, Edelsten D (1997), ‘Low cross-flow velocity microfiltration of skim milk for removal of bacterial spores’, Int Dairy J 7, 849–861. Hongvaleerat C, Cabral L M C, Dornier M, Reynes M, Ningsanond S (2008), ‘Concentration of pineapple juice by osmotic evaporation’, J Food Eng 88, 548–552. Huang R Y M, Shao P, Nawawi G, Feng X, Burns C M (2001), ‘Measurements of partition, diffusion coefficients of solvents in polymer membranes using rectangular thin-channel column inverse gas chromatography’, J Membr Sci 188, 205–218. Hughes D, Field R W (2006), ‘Crossflow filtration of washed and unwashed yeast suspensions at constant shear under nominally sub-critical conditions’, J Membr Sci 280, 89–98. Jaffrin M J (2008), ‘Dynamic shear-enhanced membrane filtration: a review of rotating disks, rotating membranes and vibrating systems’, J Membr Sci, 324, 7–25. Jaffrin M Y, Ding L H, Akoum O, Brou A (2004), ‘A hydrodynamic comparison between rotating disk and vibratory dynamic filtration systems’, J Membr Sci 242, 155–167. Jefferson B, Laine A T, Judd S J, Stephenson T (2000), ‘Membrane bioreactors and their role in wastewater reuse’, Water Sci Technol 41, 197–204. Jiao B, Cassano A, Drioli E (2004), ‘Recent advances on membrane processes for the concentration of fruit juices: a review’, J Food Eng 63, 303–324. Johnson J R, Braddock R J, Chen C S (1996), ‘Flavor losses in orange juice during ultrafiltration and subsequent evaporation’, J Food Sci 61, 540–543. Jonsson G (1985), ‘Molecular weight cut-off curves for ultrafiltration membranes of varying pore sizes’, Desalination 53, 3–10. Jonsson G (2008), ‘Tuning of the cut-off curves by dynamic ultrafiltration’, Proceedings of the International Conference on Membranes and Membrane Processes: ICOM2008, Hawaii, July 12–18. Jonsson G, Boesen C E (1975), ‘Water and solute transport through cellulose acetate reverse osmosis membranes’, Desalination 17, 145–165. Kamada T, Nakajima M, Nabetani H, Iwamoto N S (2002), ‘Availability of membrane technology for purifying and concentrationg oligosaccharides’, Eur Food Res Technol 214, 435–440. Khayet M, Godino M P, Mengual J I (2002), ‘Thermal boundary layers in sweeping gas membrane distillation processes’, AIChE J 48, 1488–1497. Koroknai B, Csanadi Z, Gubicza L, Belafi-Bako (2008), ‘Preservation of antioxidant capacity and flux enhancement in concentration of red fruit juices by membrane processes’, Desalination 228, 295–301. Kozak A, Bekassy E, Vatai G (2009), ‘Production of black-currant juice concentrate by using membrane distillation’, Desalination 241, 309–314. Kozak A, Rektor A, Vatai G (2006), ‘Integrated large-scale membrane process for producing concentrated fruit juices’, Desalination 200, 540–542. Lagana F, Barbieri G, Drioli E (2000), ‘Direct contact membrane distillation: modeling and concentration experiments’, J Membr Sci 166, 1–11. Lawson K W, Lloyd D R (1997), ‘Membrane distillation’, J Membr Sci 124, 1–25. Lazaridou A, Biliaderis C G (2007), ‘Molecular aspects of cereal beta-glucan functionality: physical properties, technological applications and physiological effects’, J Cereal Sci 46, 101–118. Li W, Li J, Chen T, Chen C (2004), ‘Study on nanofiltration for purifying fructooligosaccharides II. Extended pore model’, J Membr Sci 258, 8–15. Machado D R, Hasson D, Semiat R (2000), ‘Effects of solvent properties on permeate through nanofiltration membranes. Part II: transport model’, J Membr Sci 166, 6–69.
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200 Separation, extraction and concentration processes Makardij-Tossonian A A (2009), ‘Membrane regeneration’, World Patent WO/2009/089587. Melin T, Jefferson B, Bixio D, Thoeye C, De Wilde W, De Koning J, van der Graaf J, Wintgens T (2006), ‘Membrane bioreactor technology for wastewater treatment and reuse’, Desalination 187, 271–282. Mengual J I, Khayet M, Godino M P (2004), ‘Heat and mass transfer in vacuum membrane distillation’, Int J Heat Mass Transfer 47, 865–875. Mickols W, Koreltz M S, Moll D J, Streeter D B, Mickols W, Kobeltz M S (2008), ‘Spiral wound module assembly for e.g. in-situ on-line and real-time measurement of membrane fouling in spring wound module, has membrane envelopes and acoustic transducers which are located adjacent to permeate collection tube’, World Patent WO2008103864. Mulder M (1996), Basic principles of membrane technology, Dordrecht/Boston/London: Kluwer Academic Publishers. Nabarlatz D, Torras C, Garcia-Valls R, Montane D (2007), ‘Purification of xylooligosaccharides from almond shells by ultrafiltration’, Sep Purif Technol 53, 235–243. Nigam M O, Bansal B, Chen X D (2008), ‘Fouling and cleaning of whey protein concentrate fouled ultrafiltration membranes’, Desalination 218, 313–322. Nishizawa K, Nakajima M, Nabetani H (2000), ‘A forced flow membrane reactor for transfructosylation using ceramic membrane’, Biotechnol Bioeng 68, 92–97. Olano-Martin E, Mountzouris K C, Gibson G R, Rastall R A (2001), ‘Continuous production of pectic oligosaccharides in a membrane enzyme reactor’, J Food Sci 66, 966–971. Omprakash M, Bansal B, Cheng D X (2008), ‘Fouling and cleaning of whey protein concentrate fouled ultrafiltration membranes’, Desalination 218, 313–322. Oussedik S, Belhocine D, Grib H, Loucini H, Piron D L, Nameri N (2000), ‘Enhanced ultrafiltration of bovine serum albumin with pulsed electric field and fluidized activated alumina’, Desalination 127, 59–64. Palmieri L, Dalla Rosa M, Dall’Aglio G, Carpi G (1990), ‘Production of kiwifruit concentrate by reverse osmosis process’, Acta Horticulturae 282, 435–439. Perle M, Kimchie S, Shelef G (1995), ‘Some biochemical aspects of the anaerobic degradation of dairy wastewater’, Water Res 29, 1549–1554. Pinelo M, Jonsson G, Meyer A S (2009), ‘Membrane technology for purification of enzymatically produced oligosaccharides: molecular and operational features affecting performance’, Sep Purif Technol, 70, 1–11. Ramaswami H S, Marcotte M (2005), Food processing: principles and applications, CRC Press. Rice G, Barber A, O’Connor A, Stevens G, Kentish S (2009a), ‘Fouling of nanofiltration membranes by dairy ultrafiltration permeates’, J Membr Sci 330, 117–126. Rice G, Kentish S, O’Connor A, Barber A, Pihlajamäki A, Nyström M, Stevens G (2009b), ‘Analysis of separation and fouling behaviour during nanofiltration of dairy ultrafiltration permeates’, Desalination 236, 23–29. Sanz M L, Polemis N, Morales N, Corzo A (2005), ‘In vitro investigation into the potential prebiotic activity of potential oligosaccharides’, J Agric Food Chem 53, 2914–2921. Sarney D B, Hale C, Frankel C, Vulfson E N (2000), ‘A novel approach to the recovery of biologically active oligosaccharides from milk using a combination of enzymatic treatment and nanofiltration’, Biotechnol Bioeng 69, 461–467. Sheu M J, Wiley R C (1983), ‘Preconcentration of apple juice by reverse osmosis’ J Food Sci, 48, 422–429. Sjöman E, Mänttäri M, Nyström M, Koivikko H, Heikkilä H (2008), ‘Xylose recovery by nanofiltration from different hemicellulose hydrolyzate feeds’, J Membr Sci 310, 268–277.
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Advances in the effective application of membrane technologies 201 Swennen C M, Coyrtin B, Bruggen C, Vandecasteele C, Gibson G R, Rastall R A (2005), ‘Ultrafiltration and ethanol precipitation for isolation of arabinoxylooligosaccharides with different structures’, Carbohydr Polym 62, 283–292. Tohammadi T, Madaeni S S, Moghadam M K (2002), ‘Investigation of membrane fouling’, Desalination, 153, 155–160. Trägårdh G (1989), ‘Membrane cleaning’, Desalination, 71, 325–335. Valdés H, Romero J, Saavedra A, Plaza A, Bubnovich V (2009), ‘Concentration of noni juice by means of osmotic distillation’, J Membr Sci 330, 205–213. Vegas R, Luque S, Alvarez J R, Alonso J L, Dominguez H, Parajo R C (2006), ‘Membraneassisted processing of xylooligosaccharides-containing liquors’, J Agric Food Chem 54, 5430–5436. Vidal G, Carvalho A, Mendez R, Lema J M (2000), ‘Influence of the content in fats and proteins on the anaerobic biodegradability of dairy wastewaters’, Bioresour Technol 74, 231–239. Wheatley A (1990), Anaerobic digestion: a waste treatment technology. London and New York: Elsevier Applied Science.
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7 Electrodialytic phenomena, associated electromembrane technologies and applications in the food, beverage and nutraceutical industries L. Bazinet, A. Doyen and C. Roblet, Laval University, Canada Abstract: Electrodialysis, an electrochemical separation process with charged membranes stacked to separate ionic species from aqueous solutions and uncharged components when an electrical field is applied, is providing new membrane separation processes with numerous applications in the food, nutraceutical and beverage industries. Techniques such as electrolysis with membrane, electrodialysis with ionexchange membranes, electrodialysis with bipolar membrane and electrodialysis with filtration membrane have been used for, among other applications, the coagulation of protein, the electroreduction of the medium and/or the fractionation of several food proteins. Moreover, electrodialysis with filtration membranes has been used for the recovery of molecules with bioactive properties such as antioxidant, anticancer and antihypertensive peptides. Key words: electrodialysis, electromembranes, nutraceuticals, biomolecules, purification, separation.
7.1 Introduction Electrodialysis (ED) is one of a group of membrane-based separation technologies which are finding increasing use, in agri-food industries to concentrate, purify or modify foods. In ED, an electric field provides the driving force and porous or non-porous membranes perform the separation: electrodialysis is a combined method of dialysis and electrolysis (Shaposhnik and Kesore, 1997). ED can be performed with two main cell types: electrolysis (or electro-electrodialysis) cells for oxido-reduction reactions and multimembrane cells for dilution–concentration, water dissociation and purification applications. The electrolysis cell operates with only one membrane that © Woodhead Publishing Limited, 2010
Electrodialytic phenomena and associated electromembrane technologies 203
separates two solutions circulating in each electrode compartment. This application is based on electrode redox reactions that are electrolysis-specific properties. In multi-membrane cells only the membrane transport phenomena intervenes, whereas electrochemical reactions occurring at the electrodes do not interact with the separation process. The principles of electrode and membrane reactions, as well as the technologies associated with these electrodialytic phenomena are reviewed in this chapter. The specific applications of electrodialysis and electrolysis currently used and those under development in the food, beverage and nutraceutical industries are presented.
7.2 Principles of electrodialytic phenomena and associated membrane technologies 7.2.1 Membrane phenomena An ion-exchange membrane is made of a macromolecular material (skeleton) which carries ionizable groups such as ion-exchange resins. The membrane contains fixed ions firmly attached to the skeleton and is electronically neutralized by mobile charges of the opposite sign, called counterions. Counterions, which carry current in the membrane, are positive in the case of cation-exchange membranes (CEM), and negative for anion-exchange membranes (AEM) (Bazinet, 2005). Both of these membranes are monopolar; this means that they are permeable to only one type of ion (Gardais, 1990). The perm-selectivity of the membrane is the result of an electrostatic repulsion called Donnan exclusion (Donnan, 1911). For a cation-exchange membrane (CEM), anions are repulsed from the membrane thus allowing only cations to migrate through the membrane. 7.2.2 Electrode phenomena In membrane electrolysis, membrane selectivity, explained above, and electrolysis phenomenon are both active. In electrolysis, an external potential difference is applied to the cell and chemical reactions occur at the electrode–solution interface. The faradaic reactions, caused by the passage of a current, are characterized by electron transfer at the electrode–solution interface. The oxidation A Æ Az+ + z e– induces the loss of one or more electrons and the reduction Bn+ + n e– Æ B induces the gain of one or more electrons. The anode induces oxidation reactions, and reduction reactions occur at the cathode. This transfer always occurs at the electrode surface (Brett and Oliveira-Brett, 1994; Gardais, 1990).
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204 Separation, extraction and concentration processes
7.3 Applications of electrodialytic phenomena and associated membrane technologies Only one type of membrane technology is associated with electrode reactions (electrolysis with membrane) whereas three types are associated with membrane phenomena (electrodialysis with ion-exchange membranes, electrodialysis with bipolar membrane, and, recently, electrodialysis with filtration membranes) (Fig. 7.1). 7.3.1 Electrolysis with membrane Two main applications of electrode reactions are electrochemical coagulation (EC) and electroreduction. Both methods are exploratory, but show an interesting potential for applications in the food, beverage and nutraceutical industry (Table 7.1). An example of both applications will be developed. Milk protein electrochemical coagulation Khidirov and Merzametov (1982) used EC to precipitate milk proteins in an electrolysis cell separated by a ceramic diaphragm. The milk is poured into the anodic compartment and an electrolyte or a whey solution is poured into the cathodic compartment. Proteins coagulate on the platinum electrode by forming a highly dense white coagulum. Every type of milk could be treated for protein electrochemical coagulation without the use of rennet. Both the caseins and the whey proteins coagulate at the anode. Janson and Lewis (1994) studied the possible use of electrochemical coagulation to directly separate up to 73.8% of the total cheese whey proteins. The whey was circulated in the anodic compartment of an electrolysis cell with a non-ion-selective membrane, made of cotton, for the acidification–coagulation phase (Janson et al., 1990). Afterwards the solution was separated into the coagulum and an impoverished protein solution. This solution was then circulated in the cathodic compartment of the electrolysis cell to undergo the alkalinization phase. Enhancement of lipid stability of omega-3 enriched commercial milk Consumer demand for specific nutritional qualities is encouraging the dairy industry to develop products supplemented in omega-3 fatty acids. Although these fatty acids are known to have many health benefits, they are extremely susceptible to oxidative deterioration which causes difficulties during their storage. In the study by Haratifar (2008) an electroreduction process was performed to modify the redox state of omega-3 enriched milk. A 4-V electroreduction treatment was applied for 1 h on pasteurized omega-3 enriched milk, at room temperature. The electroreduction treatment reduced the redox potential value of omega-3 enriched milk samples quickly and decreased their dissolved oxygen concentration. The electroreduced and control samples were stored at room temperature for up to 3 weeks in the presence of fluorescent light and
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(b)
2H2O +
2e– + 2H2O H+
O2 + 4H+ + 4e– OH–
Anode
H2 +
Desalted solution
AEM CEM
AEM CEM
Cl–
– Cathode
+ Anode
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2OH–
Cl
–
Cl Na
Cl–
–
Na+ Na+
+
–
Na+
Cathode n
Feed solution (c)
Anionic/acid fraction
AEM FM
Neutral fraction
FM CEM
FM
P+/– + Anode
A
–
P–
P+ A– C+ P–
P+
A
(d)
Cationic/basic fraction
–
P–
FM CEM P+/–
Cl–
A– P+ P+ C
NaOH solution
AEM CEM BPM AEM Cl–
HCl solution CEM BPM AEM H+
H+
+
C
P–
+
–
+
Cathode
Anode
OH Na+
–
Na+ Cl–
Cl
–
Na+
n Feed solution
OH– Na+
Cl–
CEM
Cl–
–
Na+
Na+ Cathode n
Salt solution
Fig. 7.1 Electrodialytic phenomena and associated membrane configurations: (a) redox reaction; (b)–(d) membrane reactions. (a) Electrolysis with membrane; (b) electrodialysis with ion-exchange membranes; (c) electrodialysis with filtration membranes; and (d) electrodialysis with bipolar membranes. AEM, anion-exchange membrane; CEM, cation-exchange membrane; BPM, bipolar membrane; FM, filtration membrane; A–, anion; C+, cation; P–, anionic peptide; P+, cationic peptide; P+/–, neutral peptide.
Electrodialytic phenomena and associated electromembrane technologies 205
(a)
206 Separation, extraction and concentration processes in the dark and the composition in fatty acids of milk samples was measured by gas chromatography with flame ionization detection. It was shown that storage under fluorescent light involved a degradation of the fatty-acids, whereas the electroreduction treatment slowed down the oxidative degradation of electroreduced samples compared with untreated milk samples. Results of this study show that the electroreduction treatment can be a potential method of enhancing the shelf-life of products containing unsaturated fatty acids. 7.3.2 Electrodialysis with ion-exchange membranes Conventional electrodialysis or electrodialysis with ion-exchange membranes consist of a series of cation- and anion-exchange membranes arranged in an alternating pattern between an anode and a cathode to form individual cells. Under the influence of an applied potential gradient between cathode and anode, the positively charged cations migrate towards the cathode and negatively charged anions towards anode (Bazinet, 2005, Bazinet and Firdaous, 2009b). The main applications of ED and dilution–concentration in the food industry consist of the demineralization of milk, sugar and whey. In the 1990s, a new application was developed for tartaric stabilisation of wine. The use of ED in membrane bioreactors, for protein separation and acid production was studied. The main applications of conventional ED are presented in Table 7.1, and the application on tartaric stabilization of wine is described in detail here. The use of ED for tartaric stabilization of wine has been approved by the Council of Europe in 1998 and is presently used worldwide on an industrial scale. Wine contains tartaric acid (H2T), a dicarboxylic acid which dissociates in the tartrate (HT–) and bitartrate (T2–) forms. During alcoholic fermentation, potassium and tartrate ions associate and precipitate to form potassium tartrate crystals (KHT) (Gonçalves et al., 2003). As a consequence, at normal storage temperatures, an undesirable KHT precipitation occurs in wine bottles. The ED process for wine tartaric stabilization was developed by Escudier et al. (1995). The principle is the same as for the conventional ED process with a diluate compartment in which the wine to be treated circulates, under nitrogen bubbling, and a concentrate compartment where salts extracted from the wine are recovered. According to the authors, the level of wine deionization, based on the decrease of the conductivity, should be 5 to 20% to obtain tartaric stabilization of wine. One of the main advantages of this technique is the preservation of the organoleptic properties of wine during the stabilization process. 7.3.3 Electrodialysis with filtration membranes Bazinet et al. (2005) stacked filtration membranes and ion-exchange membranes in a conventional electrodialysis cell. This technology, named electrodialysis
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Technology
Substrate
Target compounds Objective
Applications and stage of development
References
Dissolved oxygen
Protection of flavour after spray-drying
Food (in development)
Inoue and Kato, 2003
Fruit juices
Dissolved oxygen
Increase in shelf-life
Food (in development)
Hekal, 1983 Swanson and Sommer, 1940
Milk
W-3 fatty acids
Protection against lipid oxidation
Functional food (in development)
Haratifar, 2008
Dissolved oxygen
Protection of flavour after spray-drying
Food (in development)
Inoue and Kato, 2003
Enhancement of probiotic bacteria growth
Functional food (in development)
Bolduc et al., 2006
Food (in development)
Mondal and Lalvani, 2003
Electrode reactions Electrolysis with Coffee membrane © Woodhead Publishing Limited, 2010
Oil
Trans fatty acid
Decrease trans fatty acid content
Water
Electroreductive compounds
Increase purity and quality Food of tap water (in development)
Crandall et al., 2001 Koseki et al., 2003 Mercier, 1999
Increase microbiological quality of water
Food (in development)
Kim et al., 2000 Thompson and Gerson, 1985
Electroreduction of whey protein
Food (in development)
Bazinet et al., 1997a
Production of protein isolate
Food (in development)
Janson and Lewis, 1994
Whey
Whey proteins
Electrodialytic phenomena and associated electromembrane technologies 207
Table 7.1 Electrodialytic phenomena and associated applications in the food, beverage and nutraceutical industries
Table 7.1 Continued Substrate
Membrane reactions Electrodialysis Maple sap with ion-exchange membranes Milk © Woodhead Publishing Limited, 2010
Target compounds Objective
Applications and stage of development
References
Calcium
Food (in development)
Bazinet et al., 2007
Food (in development) Food (industrial scale) Food (in development) Food (industrial scale)
Bolzer, 1985
Casein Minerals
Passion fruit Citric acid juice Sugar beet juice, Minerals sugar cane and molasses Soy tofu whey Minerals
Whey
Magnesium or calcium Lactic acid
To avoid sugar sand formation during maple syrup production Production of acid caseinate Demineralization of milk for further applications Deacidification of fruit juice Demineralization of sugar syrups Use of soy tofu whey as bacteria growth medium Recovery of coagulant agent Production of biological preservative agent
Minerals
Demineralization of whey for further applications
Propionic acid
Production of yeast inhibitor Whey protein separation
Whey
Food (in development) Food (in development) Food and biopharmaceutical (in development) Food (industrial scale) Food (in development) Food (in development)
Hiraoka et al., 1979 Vera Calle et al., 2002, 2003 Chaput, 1979 Ben Ounis et al., 2008 Bazinet et al., 1999b Boyaval et al., 1987 Glassner, 1992 Houldsworth, 1980 Pérez et al., 1994 Boyaval et al., 1993 Amundson et al., 1982 Slack et al., 1986; Stack et al., 1995
208 Separation, extraction and concentration processes
Technology
Electrodialysis with Apple juice bipolar membranes
Potassium
To avoid potassium bitartrate formation
Polyphenol oxidase Inhibition of enzymatic browning
Food (industrial scale)
Audinos et al., 1979, 1985 Escudier et al., 1995 Guérif, 1993
Food (in development)
Quoc et al., 2000, 2006 Tronc et al., 1997, 1998
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Chitosan
Chitosanase
Chitosanase inhibition and Nutraceuticals optimization of chitosane (in development) oligomer production
Lin Teng Shee et al., 2008
Milk
Casein
Production of isolate
Food (in development)
Bazinet et al., 1999a, 2001
Passion fruit juice
Citric acid
Deacidification of fruit juice
Food (in development)
Vera-Calle et al., 2002, 2003
Soybean
Soy proteins
Production of isolate
Food (in development)
Bazinet et al., 1996, 1997b, 1997c, 1998 Mondor et al., 2004 Skorepova and Moresoli, 2007
11S and 7S
Fractionation of 7S and 11S fractions
Food (in development)
Bazinet et al., 2000
Soy proteins
Recovery of soy proteins
Food (in development)
Bazinet et al., 1999b
Calcium or magnesium
Recovery of coagulant agent
Food (in development)
Bazinet et al., 1999b
a-lactalbumin (a-la) Production of a-la and and b-lactoglobulin b-lg enriched fractions (b-lg)
Food (in development)
Bazinet et al., 2004a, 2004b
Soy tofu whey
Whey
Phospholipids
Production of phospholipid Food and nutraceuticals Lin Teng Shee et al., 2005, (in development) 2007 and protein enriched fractions
Electrodialytic phenomena and associated electromembrane technologies 209
Wine
Technology
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Electrodialysis with filtration membranes and electromembrane filtration
Substrate
Target compounds Objective
Applications and stage of development
References
Whey protein
Prodution of isolate
Food (in development)
Bazinet et al., 2004a
Lactic acid
Production of biologic preservative agent
Food and biopharmaceutical (in development)
Norddahl, 1998 Norddahl et al., 2001
Peptide VW
Production of an ACE inhibitor peptidic fraction
Food, nutraceutical and bio-pharmaceutical (in development)
Firdaous et al., 2009 Bazinet and Firdaous, 2009a
Cranberry juice Proanthocyanidins and anthocyanins
Production of antioxidant enriched fruit juices
Food and nutraceutical (in development)
Bazinet et al., 2009 Bazinet and Firdaous, 2009b
Green tea
Catechin
Production of antioxidants Food, nutraceutical and bio-pharmaceutical (in development)
Labbé et al., 2005
Milk
as2-Casein hydrolysate (fraction 183-207)
Bioactive peptide separation
Bargeman et al., 2000, 2002
Snow crab
Protein hydrolysate Production of an anticancer Nutraceutical and peptidic fraction biopharmaceutical (in development)
Doyen et al., 2010
Whey
ACE inhibitor bioactive b-Lactoglobulin peptide separation hydrolysate (fraction 142–148)
Nutraceutical and biopharmaceutical (in development)
Bazinet et al., 2005 Poulin et al., 2006, 2007
Lactoferrin
Food (in development)
Ndiaye et al., 2010
Alfalfa white protein
Separation of lactoferrin from whey
Nutraceutical and biopharmaceutical (in development)
210 Separation, extraction and concentration processes
Table 7.1 Continued
Electrodialytic phenomena and associated electromembrane technologies 211
with filtration membranes (EDFMs), couples size-exclusion capabilities of porous membranes with the charge selectivity of electrodialysis (ED). Because no pressure is applied in the electrodialysis cell, the electric field is the only driving force (Bazinet and Firdaous, 2009b). This technology has already been tested for the separation/purification of various high-added-value bioactive molecules (Table 7.1) such as anticancer peptides and antioxidant polyphenols that will be discussed in more detail. Production of an anticancer peptidic fraction Protein hydrolysates from marine products have interesting bioactive properties such as antithrombotic, antihypertensive (Kim and Mendis, 2006), antioxidant (Amarowicz and Shahidi, 1997), anticancer (Picot et al., 2006) and antimicrobial activities (Beaulieu et al., 2010). Recently, a snow crab by-product hydrolysate, has demonstrated antibacterial inhibition properties against specific strains such as Aeromonas hydrophila, Vibrio vulnificus and Vibrio parahaemolyticus (Beaulieu et al., 2010). Further investigations on the bioactive properties of polypeptides originating from this specific snow crab by-product hydrolysate were performed by Doyen et al. (2010). They carried out an electroseparation of the mixture at pH 9, 6 and 3. An ultrafiltration membrane was placed in an electrodialysis cell on both side of hydrolysate compartment allowing the simultaneous separation and recovery of anionic and cationic peptides. The pH values of the recovery compartments were also maintained at pH 9, 6 and 3 during all the separation process. The authors concluded that a selective separation was obtained by the recovery of two anionic and three cationic peptides in the adjacent compartment. They also tested the peptidic fractions for their anti-cancer activity. The results showed that only one fraction amongst the six fractions produced demonstrated a significant inhibition of growth cancerous lines especially on A549 (lung), BT549 (breast) and PC3 (prostate) cell lines. Furthermore, no anticancer aspect was observed for the feed hydrolysate fractions before separation. Consequently, the authors concluded that the anticancer activity was obtained after purification by electrodialysis with ultrafiltration membranes. Indeed, at a low concentration of 190 mg mL–1, cellular mortality of 75–85% for the A549 cell line, of 80% for the BT549 and of 95–100% for the PC3 cell line were obtained. Although the mechanism of action is still unknown, the future perspectives are very interesting because it is possible to isolate each peptide of the fractions recovered and to test them to determine the peptide(s) sequence(s) of interest. Antioxidant enrichment of cranberry juice Cranberry fruits are a rich source of phenolic phytochemicals including phenolic acids, flavonoids and ellagic acids (Vattem et al., 2005). Many of these compounds have antioxidant activity, equal to or greater than vitamin E (Yan et al., 2002). Cranberry juice can also prevent gastric ulcers caused
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212 Separation, extraction and concentration processes by Helicobacter pylori (Shmuely et al., 2007) and displays potent anticancer activity (Neto et al., 2008). Bazinet et al. (2009) circulated in both compartments on each side of the filtration membrane a volume of cranberry juice (respectively, 450 mL and 200 mL) to concentrate a cranberry juice in its own specific compounds. The total concentrations of proanthocyanidins and anthocyanins increased by 34.8 and 52.9%, respectively in the 200 mL cranberry juice treated by EDFM with a 500 kDa filtration membrane. Moreover, a 18% increase of the antioxidant capacity was measured by means of the oxygen radical absorbance capacity (ORAC) test in the 200 mL of enriched cranberry juice. Moreover, the taste of the enriched cranberry juice was better than the non-treated juice. Based on these results the authors concluded that the production of phenolic antioxidant enriched cranberry juice could be feasible on a large scale and proposed an integrated process flow. According to this process flow, the EDFM process would be directly connected to the bottling process of cranberry juice to produce antioxidant-enriched cranberry juices by a batch process and cranberry juice with very low variation in antioxidant in a continuous process. An advantage of this technology is that compounds of interest would be directly transferred from one juice to another without the need for solvent to extract them first. 7.3.4 Electrodialysis with bipolar membranes A new type of membrane, called a bipolar membrane (BPM), appeared commercially at the end of the 1980s. Bipolar membranes carry out the dissociation of water in the presence of an electric field. These membranes are composed of three parts: an anion-exchange layer, a cation-exchange layer, and a hydrophilic transition layer at their junctions (Mani, 1991). In electrodialysis with bipolar membranes (EDBPMs), cation- and anion-exchange membranes are stacked together with bipolar membranes in an alternating series in an electrodialysis cell and allow its application to numerous products (Table 7.1). Amongst these applications, only the deacidification of fruit juices and the fractionation of 11S-7S soybean protein is developed. Deacidification of passion fruit juices The yellow passion fruit, Passiflora edulis f. flavicarpa, has an intense and special aroma and flavour, which make it a desirable ingredient in the formulation of various food products. However, because of its high acidity, only limited amounts of juice can be added as an ingredient in the formulation of various preparations (Adhikary et al., 1983; Couture and Rousseff, 1992). To overcome this Vera Calle et al. (2002, 2003) tested the deacidification of clarified passion fruit juice by electrodialysis with bipolar membranes (EDBPMs). The stack was equipped with homopolar and bipolar membranes, forming two or three compartments. The reduction of acidity was achieved by increasing the pH from 2.9 to 4.0. This pH limit was chosen to avoid
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Electrodialytic phenomena and associated electromembrane technologies 213
micro-organism growth and spoilage (Vera Calle et al., 2002). In these EDBPM configurations only anions were able to pass through the anionexchange membranes from the juice to the concentrate compartment. The net effect was the extraction of anions, mainly citrate, and their replacement by hydroxyl ions provided by the BPM. Citric acid was formed in the concentrate compartment by citrate ions extracted from juice and protons provided by the second BPM separating the concentrate compartment from the electrode compartments. The anion concentration was similar for the two processes. The inorganic ions were almost eliminated, and 62% of the citrate ions and 48% of the malate ions were removed from the fresh juice. According to the sensory properties of deacidified juices, no significant differences were observed between the deacidified juices and the fresh juice. 11S–7S fractionation The two major reserve soybean proteins, the globulins 7S or b-conglycinin (37–39% of total protein) and 11S or glycinin (31–44% of total protein) have different intrinsic properties leading to different functional properties. From results presented by Bazinet et al. (2000), it appears that during the EDBPM process, the temperature has a large effect on the selective precipitation of the soybean protein fractions, 11S and 7S. Hence, at 27 °C, the precipitation profile of the four protein fractions is situated in a pH range from 6.6 to 4.4, with no possibility of separating any of these fractions. However, at 10 °C, the 11S globulin precipitates at a higher pH than it does at 27 °C, respectively, pH 6.7 compared with 5.9, allowing the fractionation of 11S from the other fractions. Using electroacidification, it is possible to obtain a precipitate solution enriched in the 11S fraction (71.8% of 11S and 10.8% of 7S) and a supernatant solution enriched in the 7S fraction (46.6% of 7S and 4.6% of 11S). EDBPM appears to be a promising technology for fractionation of proteins, by adjusting the electroacidification and solution parameters. This technology does not denature protein (Bazinet et al., 1996), does not use any chemical acids or bases during the process, and produces isolates with a low ash content, which results in a final product of high purity.
7.4 Future trends In this chapter, an overwiew of electromembrane technologies and their associated applications has been presented. Numerous studies have already been performed on various raw matrices (milk, soya, whey, fruit juices) using different cell configurations, but new applications await discovery. Very recently electromembrane processes allowed the separation of bioactive peptides from snow-crab by-products (Doyen et al., 2010) and alfalfa protein hydrolyzates (Firdaous et al., 2009), and the enrichment of cranberry juices in antioxidant compounds (Bazinet et al., 2009). In addition, the low energy © Woodhead Publishing Limited, 2010
214 Separation, extraction and concentration processes cost and constant selectivity during separation of EDFM by absence of transmembrane pressure and, consequently, no membrane fouling, represent an improvement over the conventional separation processes used in the agrifood industries. Furthermore, many acid or alkaline solutions traditionally used in conventional processes are generated in situ by the EDBPM cell configuration. Based on these observations, it appears that electromembrane technologies present several advantages in agri-food industries, in particular, enrichment of nutraceutical products to improve their health benefits, in response to customer demands. In the future, numerous aspects will have to be explored to enhance the understanding of the technologies. Hence, studies on (1) protein/peptide interactions with filtration and ion exchange membranes (IEM) interfaces during treatment/separation of hydrolyzate solutions, and (2) the role of hydrolyzate mineral content on the modification of peptide theoretical charges and influence on their migration rates, would be of major interest. The understanding of both phenomena could lead to the development of new membrane materials and applications, and consequently, to promising ways for new products manufacturing.
7.5 References Adhikary, S.K., Harkare, W.P., Gowindan, K.P. and Nanjundaswamy, A.M. (1983) Deacidification of fruit juices by electrodialysis, Indian J. Technol., 21: 120–123. Amarowicz, R. and Shahidi, F. (1997) Antioxidant activity of peptide fractions of capelin protein hydrolysates, Food Chem., 58: 355–359. Amundson, C.H., Watanawanichakorn, S. and Hill, C.G., Jr. (1982) Production of enriched protein fractions of b-lactoglobulin and a-lactalbumin from cheese whey. J. Food Process Preserv, 6: 55–71. Audinos, R., Roson, J.P. and Jouret, C. (1979) Application de l’électrodialyse à l’élimination de certains composants du jus de raisin et du vin. Connaissance Vigne Vin, 13: 229–239. Audinos, R., Lurton, L. and Moutounet, M. (1985) Advantage of electrodialysis to produce sweetening products from grape. Sci. Aliment., 5: 619–637. Bargeman, G., Dohmen-Speelmans, M., Recio, I., Timmer, M. and Van den Horst, C. (2000) Selective isolation of cationic amino acids and peptides by electro-membrane filtration. Lait, 80: 175–185. Bargeman, G., Houwing, J., Recio, I., Koops, G.H. and Van den Horst, C. (2002) Electromembrane filtration for the selective isolation of bioactive peptides from an as2-casein hydrolysate. Biotechnol Bioeng., 80: 599–609. Bazinet, L., Cossec, C., Gaudreau, H. and Desjardins, Y. (2009) Production of a phenolic antioxidant enriched cranberry juice by an electrochemical process. J. Agric. Food Chem., 57: 10245–10251. Bazinet, L. (2005) Electrodialytic phenomena and their applications in the dairy industry: a review. CRC Critical review in Food Science and Nutrition, 45: 307–326.
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Electrodialytic phenomena and associated electromembrane technologies 215 Bazinet, L., Amiot, J., Poulin, J.-F. and Labbé, D. and Tremblay, A. (2005) Process and system for separation of organic charged compounds. World Patent 082495A1. Bazinet, L. and Firdaous, L. (2009a) Membrane processes and devices for separation of bioactive peptides. Recent Pat. Biotechnol., 3: 61–72. Bazinet, L. and Firdaous, L. (2009b) Applications of electromembrane processes to the production of nutraceuticals or functional foods. In Handbook of membrane research: properties, performance and applications. Gorley, S.V. (Editor). Chemical engineering methods and technology series, Nova Science Publishers Inc., Hauppauge, NY. Chapitre 8, pp. 291–312. Bazinet, L., Ippersiel, D. and Lamarche, F. (1999b) Recovery of magnesium and protein from soy tofu whey by electrodialytic configurations. J. Chem. Technol. Biotechnol., 74(7): 663–668. Bazinet, L., Lamarche, F. and Ippersiel, D. (1998) Comparison of chemical and bipolar membrane electrochemical acidification for precipitation of soybean proteins. J. Agric. Food. Chem., 46: 2013–2019. Bazinet, L., Lamarche, F., Boulet, M. and Amiot, J. (1997a) Combined effect of pH and temperature during electroreduction of whey proteins. J. Agric. Food Chem., 45: 101–107. Bazinet, L., Lamarche, F., Labrecque, R. and Ippersiel, D. (1997b) Effect of KCl and soya protein concentrations on the performance of bipolar membrane electro-acidification. J. Agric. Food. Chem., 45: 2419–2425. Bazinet, L., Lamarche, F., Labrecque, R. and Ippersiel, D. (1997c) Effect of number of bipolar membranes and temperature on the performance of bipolar membrane electroacidification. J. Agric. Food. Chem., 45: 3788–3794. Bazinet, L., Gaudreau, H., Lavigne, D. and Martin, N. (2007) Partial demineralization of maple sap by electrodialysis: impact on syrup sensory and physicochemical quality. J. Sci. Food Agric., 87: 1691–1698. Bazinet, L., Ippersiel, D., Gendron, C., Mahdavi, B., Amiot, J. and Lamarche, F. (2001) Effect of added salt and increase in ionic strength on skim milk electroacidification performances. J. Dairy Res., 68: 237–250. Bazinet, L., Ippersiel, D., Labrecque, R. and Lamarche, F. (2000) Effect of temperature on the separation of soybean 11S and 7S protein fractions during bipolar membrane electroacidification. Biotechnol. Prog., 16: 292–295. Bazinet, L., Ippersiel, D. and Mahdavi, B. (2004a) Effect of conductivity adjustment on the separation of whey protein by bipolar membrane electroacidification. J. Agric. Food Chem., 52: 1980–1984. Bazinet, L., Ippersiel, D. and Mahdavi, B. (2004b) Fractionation of whey protein by bipolar membrane electroacidification. Innov. Food Sci. Emerg. Technol., 5: 17–25. Bazinet, L., Lamarche, F., Ippersiel, D. and Amiot, J. (1999a) Bipolar membrane electro-acidification to produce bovine milk casein isolate. J. Agric. Food Chem., 47: 5291–5296. Bazinet, L., Lamarche, F., Labrecque, R., Toupin, R., Boulet, M. and Ippersiel, D. (1996) Systematic study on the preparation of a food grade soyabean protein. Rep. Can. Electr. Assoc., 9326, U 987, Research and Development; Montréal. Beaulieu, L., Thibodeau, J., Desbiens, M., Saint-Louis, R. Zatylny, C. and Thibaut, S. (2010) Evidence of antimicrobial activities in peptide fractions originating from snow crab (Chionoecetes opilio) by-products. Probiot. Antimicrob. Prot., DOI 10.1007/ s12602-010-9043-6, in press. Ben Ounis, W., Champagne, C. P., Makhlouf, J. and Bazinet, L. (2008) Utilization of tofu whey pre-treated by electromembrane process as a growth medium for Lactobacillus plantarum LB17. Desalination, 229: 192–203. Bolduc, M.P., Bazinet, L., Lessard, J., Chapuzet, J.M. and Vuillemard, J.C. (2006) Electrochemical modification of the redox potential of pasteurized milk and its evolution during storage. J. Agric. Food Chem., 54: 4651–4657.
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216 Separation, extraction and concentration processes Bolzer, R. (1985) Installation and process for the preparation of acid caseinates. United States Patent 4 559 119. Boyaval, P., Corre, C. and Terre, S. (1987) Continuous lactic acid fermentation with concentrated product recovery by ultrafiltration and electrodialysis. Biotech. Lett., 9: 207–212. Boyaval, P., Seta, J. and Gavach, C. (1993) Concentrated propionic acid production by electrodialysis. Enzyme Microb. Technol.; 15: 683–686. Brett, C.M.A. and Oliveira-Brett, A.M. (1994) Fundamentals of kinetics and mechanism of electrode reactions. In Electrochemistry: principles, methods, and applications; Oxford University Press: New York. Chaput, M. (1979) Le lactose, extraction et hydrolyse et déminéralisation du lactosérum. Rev. Lait. Française, 372: 23–26. Couture, R. and Rouseff, R. (1992) Debittering and deacidifying sour orange (Citrus aurantium) juice using neutral and anion exchange resins. J. Food Sci., 57: 380– 384. Crandall, J.J., Mantz, B.W. and Martz, E.C. (2001) Method for generating oxygenated water, United States Patent 6 284 293 B1. Donnan, F.G. (1911) Theory of membrane equilibrium and membrane potential in the presence of non-dialysing electrolytes: A contribution to physical-chemical physiology. Z. Elektrochem. Angewandte Phys. Chem., 17: 572–581. Doyen, A., Beaulieu, L., Saucier, L., Pouliot, Y. and Bazinet, L. (2010) Peptides from a snow crab by-products hydrolysate demonstrated anticancer properties in vitro after simultaneous separation by electrodialysis with ultrafiltration membranes. Submitted to Sep. Pur. Technol. Escudier, J., Saint-Pierre, B., Batlle, J. and Moutounet, M. (1995) Automatically controlled tartric stabilisation of wine by membrane electrodialysis which reduces its conductivity to the desired level. World Patent 9506110-A1. Firdaous, L., Dhulster, P., Amiot, J., Gaudreau, A., Lecouturier, D., Kapel, R., Lutin, F., Vézina, L.-P. and Bazinet, L. (2009) Concentration and selective separation of bioactive peptides from an alfalfa white protein hydrolysate by electrodialysis with ultrafiltration membranes. J. Membr. Sci., 329: 60–67. Gardais, D. (1990) Les procédés électriques de traitement des rejets industriels. In Environnement et Electricité. Electra, Doppee diffusion, Avon, pp. 200–310. Glassner, D. (1992) ED Applications in biotechnology. In Proceedings 10th annual membrane technology planning conference, Business Communications, Norwalk, pp. 158–165. Gonçalves, F., Fernandes, C., Cameira Dos Santos, P. and De Pinho, M.N. (2003) Wine tartaric stabilization by electrodialysis and its assessment by the saturation temperature. J. Food Eng., 59: 229–235. Guérif, G. (1993) Electrodialysis applied to tartaric stabilisation of wines. Rev. Oenol. Techn. Vitic. Œnol., 69: 39–42. Haratifar, S. (2008) The stability of electro-reduced milk lipids. Master of Science Thesis, Faculty of Agricultural and Food Sciences, Université Laval., Québec, Canada, 94p. Hekal, I.M. (1983) Process for the preservation of color and flavour in liquid containing comestibles. United States Patent 4.374.714. Hiraoka, Y., Itoh, K. and Taneya, S. (1979) Demineralization of cheese whey and skimmed milk by electrodialysis with ion exchange membranes. Milchwissenschaft, 34: 397–400. Houldsworth, D.W. (1980) Demineralization of whey by means of ion exchange and electrodialysis. J. Soc. Dairy Technol., 33: 45–51. Inoue, T.T. and Kato, W.K. (2003). Powdery drinks and process for producing the same. International Patent WO03053153. Janson, H.V. and Lewis, M.J. (1994) Electrochemical coagulation of whey protein. J. Soc. Dairy Technol., 47: 87–90. © Woodhead Publishing Limited, 2010
Electrodialytic phenomena and associated electromembrane technologies 217 Janson, H.V., Kiis, A.A., Alekseev, N.G. and Kirm, A.A. (1990) USSR Patent 1 570 694, 1990. Khidirov, S.S. and Merzametov, M.M. (1982) Electrocoagulation of milk proteins, brief communications, Vol.1, Book 2, XXI International Dairy Congress, Moscow, 12–16 July 1982. Kim, C., Hung, Y.C. and Bracket, R.E. (2000) Roles of oxidation–reduction potential in electrolyzed oxidizing and chemically modified water for the inactivation of foodrelated pathogens. J. Food Prot., 63: 19–24. Kim, S.-K. and Mendis, E. (2006) Bioactive compounds from marine processing byproducts. A review. Food Res. Int., 39: 383–393. Koseki, M., Nakagawa, A., Tanaka, Y., Noguchi, H. and Omochi, T. (2003) Sensory evaluation of taste of alkali-ion water and bottled mineral waters. J. Food Sci., 68: 354–358. Labbé, D., Araya-Farias, M., Tremblay, A. and Bazinet, L. (2005) Electromigration feasability of green tea catechins. J. Membr. Sci., 254: 101–109. Lin Teng Shee, F., Angers, P. and Bazinet, L. (2005) Precipitation of cheddar cheese whey lipids by electrochemical acidification. J. Agric. Food Chem., 53: 5635–5639. Lin Teng Shee, F., Angers, P. and Bazinet, L. (2007) Delipidation of a whey protein concentrate by electroacidification with bipolar membranes (BMEA). J. Agric. Food Chem., 55: 3985–3989. Lin Teng Shee, F., Arul, J., Brunet, S. and Bazinet, L. (2008) Performing a three-step process for conversion of chitosan to its oligomers using an unique bipolar membrane electrodialysis system. J. Agric. Food Chem., 56: 10019–10026. Mani, K.N. (1991) Electrodialysis water splitting technology. J. Membr. Sci., 58: 117–138. Mercier, D. (1999) Electrochemical treatment method and device for softening water. United States Patent 5.897.765. Mondal, K. and Lalvani, S.B. (2003) Electrochemical hydrogenation of canola oil using a hydrogen transfer agent. J. Amer. Oil Chem. Soc., 80: 1135–1141. Mondor, M., Ippersiel, D., Lamarche, F. and Boye, J.I. (2004) Production of soy protein concentrates using a combination of electroacidification and ultrafiltration. J. Agric. Food Chem., 52: 6991–6996. Ndiaye, N., Pouliot, Y., Saucier, L., Beaulieu, L. and Bazinet, L. (2010) Electroseparation of bovine lactoferrin from model and whey solutions. Sep. Pur. Technol., 74: 93–99, DOI 10.1016/j.seppur.2010.05.011. Neto, C.C., Amoroso, J.W. and Liberty, A.M. (2008) Anticancer activities of cranberry phytochemicals: an update. Mol. Nutr. Food Res., 52: S18–S27. Norddahl, B. (1998) Fermentative production and isolation of lactic acid. World Patent 9 828 433. Norddahl, B., Eriksen, S. and Pedersen, F.M. (2001) Method for producing lactic acid. World Patent 01 92555-A1. Pérez, A., Andrés, L.J., Alvarez, R., Coca, J. and Hill, C.G. (1994) Electrodialysis of whey permeates and retentates obtained by ultrafiltration. J. Food Proc. Eng., 17: 177–190. Picot, L., Bordenave, S., Didelot, S., Fruitier-Arnaudin, I., Sannier, F., Thorkelsson, G., Bergé, J.P., Guérard, F., Chabeaud, A. and Piot. J.M. (2006) Antiproliferative activity of fish protein hydrolysates on human breast cancer cell lines. Process Biochem., 41: 1217–1222. Poulin, J.-F., Amiot, J. and Bazinet, L. (2006) Simultaneous separation of acid and basic bioactive peptides by electrodialysis with ultrafiltration membrane. J. Biotechnol., 123: 314–328. Poulin, J.-F., Amiot, J. and Bazinet, L. (2007) Improved peptide fractionation by electrodialysis with ultrafiltration membrane: influence of ultrafiltration membrane stacking and electrical field strength. J. Membr. Sci., 299: 83–90.
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218 Separation, extraction and concentration processes Quoc, A., Lamarche, F. and Makhlouf, J. (2000) Acceleration of pH variation in cloudy apple juice using electrodialysis with bipolar membranes. J. Agric. Food Chem., 48: 2160–2166. Quoc, A.L., Mondor, M., Lamarche, F., Ippersiel, D., Bazinet, L. and Makhlouf, J. (2006) Effect of a combination of electrodialysis with bipolar membranes and mild heat treatment on the browning and opalescence stability of cloudy apple juice. Food Res. Int., 39: 755–760. Shaposhnik, V.A. and Kesore, K. (1997) An early history of electrodialysis with permselective membranes. J. Membr. Sci., 136: 35–39. Shmuely, H., Yahav, J., Samra, Z., Chodick, G., Koren, R., Niv, Y. and Ofek I. (2007) Effect of cranberry juice on eradication of Helicobacter pylori in patients treated with antibiotics and a proton pump inhibitor. Mol. Nutr. Food Res., 51: 746–751. Skorepova, J. and Moresoli, C. (2007) Carbohydrate and mineral removal during the production of low-phytate soy protein isolate by combined electroacidification and high shear tangential flow ultrafiltration. J. Agric. Food Chem., 55: 5645–5652. Slack, A.W., Amundson, C.H. and Hill, C.G. (1986) Production of enriched b-lactoglobulin and a-lactalbumin whey protein fractions. J. Food Proc. Pres., 10: 19–30. Stack, F.M., Hennessy, M., Mulvihill, D. and O’Kennedy, B.T. (1995) World Patent 9534216-C1. Swanson, A.M. and Sommer, H.H. (1940) Oxidized flavor in milk: II. The relation of oxidation–reduction potentials to its development. J. Dairy Sci., 23: 597–614. Thompson, B.G. and Gerson, D.F. (1985) Electrochemical control of redox potential in batch cultures of Escherichia coli. Biotechnol. Bioeng., 27: 1512–1515. Tronc, J.-S., Lamarche, F. and Makhlouf, J. (1997) Enzymatic browning inhibition in cloudy apple juice by electrodialysis. J. Food Sci. 62: 75–78. Tronc, J.-S., Lamarche, F. and Makhlouf, J. (1998) Effect of pH variation by electrodialysis on the inhibition of enzymatic browning in cloudy apple juice. J. Agric. Food Chem., 46: 829–833. Vattem, D.A., Ghaedian, R. and Shetty, K. (2005) Enhancing health benefits of berries through phenolic antioxidant enrichment: focus on cranberry. Asia Pac. J. Clin. Nutr., 14: 120–130. Vera Calle, E., Dornier, M., Sandeaux, J., Pourcelly, G. (2002) Deacidification of the clarified passion fruit juice using different electrodialysis configurations. Desalination, 149: 357–361. Vera Calle, E., Ruales, J., Sandeaux, J., Dornier, M., Persin, F., Reynes, M. and Pourcelly, G. (2003) Comparison of different methods for deacidification of clarified passion fruit juice. J. Food Eng., 59: 361–367. Yan, X., Murphy, B.T., Hammond, G.B., Vinson, J.A. and Neto, C.C. (2002) Antioxidant activities and antitumor screening of extracts from cranberry fruit (Vaccinium macrocarpon). J. Agric. Food Chem., 50: 5844–5849.
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Principles of pervaporation for the recovery of aroma compounds 219
8 Principles of pervaporation for the recovery of aroma compounds and applications in the food and beverage industries S. Sahin, Middle East Technical University, Turkey Abstract: The principles and transport mechanism of pervaporation, a membrane separation process in which the components from a liquid mixture permeate selectively through a dense membrane, are described. Information on membrane materials is given and studies on the recovery of aroma compounds by pervaporation, one of its most important applications in the food industry, are reviewed. Sources of further information are listed and future trends explored. Key words: aroma compounds, food, pervaporation, beverages.
8.1 Introduction Processing, especially at high temperatures, may cause considerable physical and/or chemical changes in the aroma compounds of foods. Physical losses may take place during evaporation and chemical changes of aroma compounds may occur owing to heat-induced oxidation or Maillard reactions. For example, in the fruit juice industry, changes in aroma compounds are inevitable since pasteurization of fruit juice is required in order to increase the shelf life of the product. In addition, aroma compounds are highly volatile so that most of them are lost by evaporation during the production of concentrated juice. The change/loss of aroma compounds may affect the final product quality and consumer’s acceptance. Therefore, the aroma compounds must be recovered either from the loss stream or removed before the raw material is subjected to heat treatment (aroma stripping) and added back to the concentrated juice. There are several methods for the extraction that minimize chemical changes and physical losses of aroma compounds. These methods can be
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220 Separation, extraction and concentration processes categorized as: vapor–liquid separations based on distillation/evaporation or partial condensation, separations based on gas injection adsorption and supercritical-fluid extraction. However, these separation techniques have some limitations. Pervaporation, which is a membrane process, is a promising alternative method for aroma recovery because of it has advantages over conventional aroma recovery processes such as high selectivity, low energy consumption, moderate operating temperatures, physical separation mechanism and no additive requirement (Lipnizki et al., 2002a). Compared with other membrane processes like ultrafiltration and reverse osmosis, pervaporation exhibits low fluxes. However, selectivities can be quite high. Therefore, this method is especially suitable for the recovery of highly diluted species. This method can be performed on the juice before the evaporation or pasteurization step thus preventing deterioration or loss of the heat-sensitive aroma compounds during the heat treatment. In addition to aroma recovery, this method can be used for dealcoholization of wine or beer, recovery of high-value-added components or removal of organic pollutants from waste streams, and removal of inhibitors from fermentation broths. It can also be used as an analytical separation technique.
8.2 Principles of pervaporation Pervaporation is a membrane technique in which the components from a liquid mixture are separated by means of partial vaporization through a non-porous permselective membrane. The term ‘pervaporation’ is used in order to emphasize the fact that the permeate undergoes a phase change from liquid to vapor during its transport through the barrier. The pervaporation term was first introduced by Kober (1917). In his study, it was observed that a liquid in a collodion (cellulose nitrate) bag that was suspended in the air, evaporated, although the bag was tightly closed. Kober realized that this phenomenon could be used for separation of liquid mixtures. In pervaporation, a membrane separates an upstream mixture in the liquid state from downstream permeates in the gaseous state. The downstream side is maintained at a reduced pressure to ensure the gaseous state. The liquid feed mixture is circulated in contact with the membrane. The permeate in gaseous state is collected in the liquid state after condensation on a cooled wall. A schematic diagram of a typical experimental set-up for pervaporation is shown in Fig. 8.1. The feed solution is kept in a reservoir inside the temperature-controlled water bath and it is pumped to the upper part of the pervaporation cell. The vacuum is applied to the lower compartment of the cell by use of a vacuum pump in order to provide a pressure gradient across the membrane. The vacuum pressure in the system is controlled by a vacuum controller. The permeate is collected in cold traps by condensation, mostly
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Principles of pervaporation for the recovery of aroma compounds 221 Thermocouple
TC Trap 1
Trap 3
Water bath
PC
Trap 2 Peristaltic Pervaporation pump cell
Vacuum pump
Fig. 8.1 Schematic diagram of a typical pervaporation experimental set-up; TC, temperature control; PC, pressure control (reprinted from Isci et al., 2006, with permission from Elsevier).
with liquid nitrogen at –196 °C. There are usually three traps in the set-up; the first two are connected in parallel in order to conduct the experiments continuously and to maintain the steady-state condition during data collection. The third one is connected in series to the parallel traps as a safety trap in order to ensure that practically no permeate reaches the vacuum pump. The components from the liquid feed mixture permeate selectively through a dense membrane driven by a chemical potential gradient obtained by partial pressure reduction on the permeate side. The affinity between the permeant and the polymer material that constitutes the membrane as well as its mobility through the membrane matrix are important for the transport of the compound. Pervaporation is a complex process involving simultaneous heat and mass transfer. Latent heat is required to pervaporate the liquid. The gradient in partial vapor pressure between the feed and the permeate side of the membrane is maintained by reducing the partial vapor pressure in the permeate side. The heat necessary for the evaporation of the permeate has to be transported through the membrane and this transport of energy is coupled to the transport of mass. The evaporation enthalpy is taken from the sensible heat of the liquid feed mixture, leading to a reduction in feed side temperature.
8.3 Transport mechanism in pervaporation for the recovery of aroma compounds Although a phase change from liquid to vapor takes place, this phase change never occurs directly but occurs via a membrane–solute interaction © Woodhead Publishing Limited, 2010
222 Separation, extraction and concentration processes in pervaporation. Therefore, the separation principle is not based on the vapor–liquid equilibrium. Transport through the nonporous membrane is only by diffusion. The solubility and the diffusivities of the constituents in the membrane are important for separation by pervaporation. Therefore, solute transport within the membrane is not directly affected by the external operating conditions. Ideally, the membrane alone determines the solute transport and the characteristics of the separation and so it is important to create the optimum mass transport conditions to and from the membrane (Schafer et al., 2006). The driving force in pervaporation is the gradient of chemical potential of a solute across the selective membrane (Schafer and Crespo, 2002). Thus, the mass transport in a pervaporation membrane can be described as a function of the chemical potential gradient:
Ji = f (D mi)
[8.1]
where Ji and mi are the mass flux and chemical potential of solute i, respectively. The chemical potential can be stated as a function of the state variables; pressure, composition and temperature:
m = f (P, C, T)
[8.2]
Therefore, the differences in the permeate pressure, feed composition and temperature cause differences in chemical potential. For this reason, there are three methods in pervaporation to establish a suitable chemical potential gradient in the membrane as driving force for the mass transport: (a) Vacuum pervaporation: a hydrostatic pressure difference between the feed and the permeate mixture is established by introducing a vacuum on the permeate side. This method is applicable if the volume of permeating vapor is relatively small or the permeate side pressure is not too low. Otherwise, vacuum pumps of extremely large capacities are required and the pumps consume too much energy (Brüschke, 2006). (b) Thermopervaporation: a chemical potential gradient in the membrane is established by providing a temperature difference between the feed and the permeate side (Strathmann and McDonogh, 1993). (c) Sweeping-gas pervaporation: an inert gas stream on the permeate side of the membrane can be used to remove the component and thus establish a concentration gradient driving force. Vacuum pervaporation is the most commonly used method because a chemical potential gradient is readily obtained by reducing partial pressure on the permeate side. Partial pressures of the permeant components are usually lower than their corresponding saturation pressures. Therefore, they are removed as vapor. The principle of solute transport across the membrane in vacuum pervaporation is shown in Fig. 8.2. A nonporous membrane separates a liquid feed, which is usually close to atmospheric pressure, from the downstream compartment where a vacuum is applied. The solutes (denoted by solid circles)
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Principles of pervaporation for the recovery of aroma compounds 223 Membrane
Liquid
Membrane
Liquid
Vapor
(a)
Vapor
(b) Membrane
Liquid
Vapor
(c)
Fig. 8.2 The principle of solute transport across the membrane in pervaporation. For explanation of panels see text.
are first absorbed by the membrane surface when they are in contact with the membrane owing to the interactions between the solutes and polymer (Fig. 8.2a). When the chemical potential of these components in the feed and in the selective layer of the membrane are equal, a thermodynamic equilibrium is reached. The solutes that are absorbed by the membrane surface create a chemical potential difference across the membrane, which causes a diffusive net flux of solute through the membrane polymer (Fig. 8.2b). Ideally, all the solute particles that have diffused through the membrane are desorbed suddenly and removed by vacuum applied on the downstream side of the membrane. As a result, the solute concentration on the membrane downstream surface remains practically zero, and a maximum concentration gradient between the two membrane surfaces is maintained. If the vacuum is not low enough to desorb all the solutes reaching the membrane downstream surface, the concentration of the solute at the bottom of the membrane is not zero and, consequently, the concentration gradient decreases, as the net flux across the membrane decreases. For the ideal diffusion of solute i across the membrane toward the membrane downstream surface, Fick’s First law applies:
dC J i = Di ÈÍ i ˘˙ Î dz ˚
[8.3]
where Ji is the flux of the permeant i, Ci is the concentration of the permeant i in the polymer membrane, z is the position in the membrane, measured
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224 Separation, extraction and concentration processes from the feed side of the membrane and Di is the diffusion coefficient of the permeant i in the membrane. The diffusion coefficient of a specific component through a polymer can be assumed to be independent of the solute concentration. Then, integration across a membrane, with a thickness Z, yields:
J = Di
(Cimem,f – Cimem,p ) Z
[8.4]
where Cimem,f and Cimem,p are the concentrations of the permeant i inside the membrane at its feed and permeate sides, respectively. If the vacuum applied on the permeate side of the membrane is strong enough, we can assume that all of the solute particles that diffuse through the membrane are removed instantaneously. Then Cimem,p can be taken as zero and equation [8.4] reduces to:
J i = Di
(Cimem,f ) Z
[8.5]
Transport, across the membrane, is generally described by the ‘solutiondiffusion mechanism’ (Strathmann and McDonogh, 1993). Assuming constant solute feed concentration and rapid removal of solute on the membrane downstream, the selective transport of the component ‘i’ from the bulk feed to the downstream compartment can ideally be described by the following steps as: 1. equilibrium partitioning (sorption) of the component i from the liquid phase at the feed solution-membrane upstream surface interphase; 2. diffusion of the absorbed component i through the membrane polymer matrix toward the downstream surface of the membrane; 3. equilibrium partitioning (desorption) of the component i from the membrane downstream surface to the permeate vapor. For the equilibrium partitioning in steps 1 and 3, the sorption (partitioning) coefficient of i in the homogeneous membrane polymer Si can be expressed as: Si =
Cimem,f Cil,bulk
[8.6]
where Cil, bulk is the concentration of component i in the bulk liquid. Combination of equations [8.5] and [8.6] yields the transport model commonly applied in pervaporation:
J i = Si Di
(Cil,bulk ) (C l,bulk ) = Pi i Z Z
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[8.7]
Principles of pervaporation for the recovery of aroma compounds 225
where Pi is the permeability of i in the membrane which is the product of the solubility and diffusivity of the solute i in the membrane. It can be concluded from equation [8.7] that, for an ideal case, the flux of a specific component is directly related to the concentration of the component in the bulk feed. In addition, the permeability of the solute can be calculated from the slope of the flux versus feed concentration graph, if the thickness of the membrane is known. The driving force can also be expressed as a partial pressure gradient of the solute i across the membrane. Then, equation [8.7] becomes:
ÈP f – Pip ˘ DP J i = Si Di ÈÍ i ˘˙ = Pi¢ Í i ˙ ÎZ ˚ Î Z ˚
[8.8]
where Pi¢ is the permeability coefficient (kg m–1 s–1 Pa–1), and Pif and Pip are partial pressures of compound i in the feed and permeate sides, respectively. The widely accepted parameters used to describe the performance of the pervaporation process are permeate flux (Ji), selectivity or separation factor (ai/j) and enrichment factor (bi). The permeate flux Ji is defined as the permeate flow rate per unit of membrane area for a given membrane thickness. The separation factor ai/j is a measure of the selectivity and it indicates the preferential permeation of component i compared with component j. It is calculated from the concentrations of the compounds to be separated, i and j, in the feed and the permeate, respectively:
a i /j =
(Cip /Cil,bulk )
(C pj /C l,bulk ) j
[8.9]
The affinity between the permeant and the polymer material that constitutes the membrane and its mobility through the membrane matrix which are responsible for the transport are important parameters for selectivity. Since different species permeate through the membrane at different rates, a substance at a low concentration in the feed stream can be highly enriched in permeate. The enrichment factor bi represents the capacity of the membrane for concentrating component i and is defined as:
bi =
Cip C l,bulk j
[8.10]
If perfect mixing of the feed and sudden removal of the solutes leaving the membrane downstream face can not be provided, boundary layers develop on both sides of the membrane, affecting the solute transport to and away from the respective membrane surface. If the flux of solute i through the membrane is higher than that through the liquid phase toward the membrane, © Woodhead Publishing Limited, 2010
226 Separation, extraction and concentration processes the solute i is depleted in the liquid phase over the membrane upstream surface, resulting in a liquid solute concentration lower than that in the bulk feed. Because the solute concentration in the liquid at the membrane upstream surface determines the partitioning of the solute into the membrane, the concentration of i in the membrane upstream surface is lower with respect to bulk concentration. Therefore, the concentration gradient across the membrane decreases as does the overall flux. This phenomenon is known as ‘concentration polarization’ and affects the fluxes of compounds of high sorption coefficient, even under turbulent hydrodynamic conditions over the membrane. A similar phenomenon can also be found on the membrane downstream face (Schafer and Crespo, 2002). The concentration polarization strongly depends on the feed flow velocity and on the hydraulic diameter. An increase in the feed flow velocity and a decrease in the hydraulic diameter, decrease the effect of concentration polarization (Borjesson et al., 1996). As long as the diffusive flux is sufficiently high compared with the solute flux across the membrane, the boundary layer does not represent an additional transport resistance and the solute concentration at the membrane surface is equal to that in the bulk (Schafer and Crespo, 2007). However, if the diffusive flux across the liquid boundary layer is lower than the maximum achievable transmembrane flux at the respective feed bulk concentration, the boundary layer detrimentally affects the overall transport of solute from the feed bulk to the permeate side of the pervaporation membrane because the membrane surface concentration of solute i is lower than that of the bulk. Overall mass transfer coefficient kov,i can be expressed as the sum of mass transfer resistances:
Ê 1 ˆ Ê 1 ˆ Ê 1 ˆ Ê 1 ˆ ÁË k ˜¯ = ÁË k ˜¯ + ÁË k ˜¯ + ÁË k ˜¯ ov,i bl,i m,i p,i
[8.11]
where kbl,i is the mass transfer coefficient for the boundary layer, km,i is the mass transfer coefficient for the membrane, and kp,i is the mass transfer coefficient for the permeate side. Although in most instances, it is hard to achieve an ideal case, the mass transfer resistance in the feed boundary layer can be neglected compared with the overall mass transfer resistance if the feed flow velocity is maintained sufficiently high during pervaporation (Karlsson and Tragardh, 1993). Moreover, if a good vacuum is applied on the permeate side of the membrane, everything that diffuses through the membrane is instantaneously pumped away and therefore, mass transfer resistance in the permeate side can also be neglected. Then, the only resistance to mass transfer is within the membrane polymer and equation [8.11] simplifies to:
Ê 1 ˆ Ê 1 ˆ ÁË k ˜¯ = ÁË k ˜¯ ov,i m,i
[8.12]
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Principles of pervaporation for the recovery of aroma compounds 227
Therefore, overall mass flux of component i from the bulk of the feed to the bulk of the permeate can be described by the following equation:
J i = kov,i (Cil,bulk – Cip ) = km,i (Cil,bulk – Cip )
[8.13]
p
Another phenomenon affecting pervaporation performance is temperature polarization. Owing to the phase change from the feed to the permeate side, the membrane acts as a heat sink and a temperature difference exists between the bulk of the feed and the feed side of the membrane. The resulting reduction in the driving force and hence in performance is called ‘temperature polarization’ and this effect is unique for pervaporation processes (Brüschke, 2006). Flux changes linearly with change in concentration but exponentially with change in temperature which means that temperature polarization is more important than concentration polarization. Therefore, the development of high flux pervaporation membranes requires the development of modules in which temperature polarization is effectively reduced. Concentration polarization does not change significantly with changes in composition of the feed. However, the effect of temperature polarization is more important if the concentration of the component to be removed, and thus the flux, is high.
8.4 Selection of membranes for pervaporation in the recovery of aroma compounds A membrane can be simply defined as a permselective barrier made up of polymers between the feed and permeate. Polymers are very large molecules (about 1000 to 100 000 times larger than water molecule) which are manufactured by chemical processes. Polymers gain different characteristics during their design. Generally, there are two types of polymers: glassy polymers and rubbery polymers or elastomers. Their ground state at room temperature determines their distinction (Koops and Smolders, 1991). Glassy polymers have a glass transition temperature above the room temperature and can be divided into three groups: crystalline, semi-crystalline and amorphous polymers. The presence of crystallites has an important affect on all kinds of polymers in terms of tensile strength, elasticity, impact strength, solubility and diffusivity in the polymer. Generally, glassy polymers are hydrophilic (Böddeker and Bengtson, 1991). They preferentially permit the permeation of water owing to the presence of groups in the polymer chain that are able to interact with water molecules. They are used mainly for removal of water from organic solvents and solvent mixtures with an emphasis on azeotropic mixtures. Membranes for the removal of small alcohol molecules like methanol or ethanol are also hydrophilic. Some examples of glassy polymers are cellulose acetate, polysulfone, and polyvinylalcohol.
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228 Separation, extraction and concentration processes Polymers with a glass transition temperature below room temperature are classified as rubbers or elastomers. Elastomers normally present higher chain mobility than glassy polymers and contain rather small, non-polar side groups. Elastomers are flexible and mostly hydrophobic polymers. They preferentially permeate organic substances owing to the lack of strong intermolecular forces, such as hydrogen bonding or dipole–dipole interactions. Since they absorb organic molecules with relatively high fluxes, elastomers are the most suitable polymers for the separation of organic molecules from aqueous solutions (Koops and Smolders, 1991) or from organic–organic mixtures. Some of the examples of hydrophobic polymers are polydimethylsiloxane (PDMS), polyether block polyamide (PEBA), polypropylene (PP), and polyvinylidenefluoride (PVDF) (Koops and Smolders, 1991). Glassy and rubbery membranes differ strongly in terms of flexibility of their polymeric structure. Within the rigid polymeric structure of a glassy membrane, the diffusivity of components is strongly related to their molecular volume, whereas in the flexible region of a rubbery membrane, solute–polymer interactions are much more important than the diffusivities of components. This is why the selectivity of glassy membranes is more diffusion controlled, whereas the selectivity of rubbery membranes is more sorption controlled (Schafer and Crespo, 2002). In pervaporation, nonporous membranes which provide selectivity based on solute–membrane interactions rather than on volatilities are used. The membrane type used in the pervaporation experiments determines the selectivity towards the molecules. Therefore, membrane selection is one of the most important issues in pervaporation. The membrane should have high selectivity and flux for the components to be separated. In addition, it must have chemical, mechanical and thermal stability (Schafer and Crespo, 1997). In pervaporation, the feed side of the membrane is highly swollen in contact with the liquid whereas the permeate side is dry and virtually nonswollen. Thus, a high gradient of swelling exists over the separating layer of the membrane, demanding additional resistance and stability. Pervaporation can be classified as hydrophilic and organophilic pervaporation (Lipnizki et al., 1999). In hydrophilic pervaporation, the target compound water is separated from an aqueous–organic mixture whereas in organophilic pervaporation, the target organic compounds are separated from an aqueous–organic mixture (hydrophobic pervaporation) or from an organic–organic mixture by being preferentially permeated through the membrane (target organophilic pervaporation). Generally, in membrane production for pervaporation processes, composite membrane structure is preferred. Composite membranes are composed of three layers. The first layer is the dense layer responsible for the separation. This layer has to be as thin as possible since the transport through the membrane is by diffusion (0.5–2 mm). The second layer is the porous support layer that has an asymmetric pore structure. It has a thickness of 70–100 mm. Structural polymers with high resistance against chemical attack and good thermal and
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Principles of pervaporation for the recovery of aroma compounds 229
mechanical properties such as polyacrylonitrile, polyetherimide, polysulfone, PVDF form the porous support. The final layer, having a thickness of about 100 mm is a woven or nonwoven textile fabric. Polyester, polyethylene, polyphenylene sulfide and similar fibers are used for the textile carrier layer (Brüschke, 2006). The dense separating layer of hydrophilic membranes is made from various polymers that have high affinity towards water. These polymers contain ions, oxygen functions such as hydroxyl-, ester, or carboxylic moieties or nitrogen as imino- or imido- groups. They must be crosslinked in order to render them insoluble after the coating process. For organophilic membranes, the dense separating layer is formed by crosslinked silicones. The most commonly used organophilic membranes are polydimethylsiloxane (PDMS) or polymethyl octyl siloxane (POMS). The monomer of this silicon rubber is: CH3 Si
CH3
O
[8.14] n
POMS is a modified silicon rubber in which one of the methyl groups of the monomer has been replaced by an octyl group. PDMS or POMS membranes with various porous support layers are commercially available. Zeolite-filled silicon membranes have been developed to increase the selectivity and flux. PEBA formulations have elastic segments consisting of thermoplastic polyamides made flexible by elastomeric polyether links. The modification of PDMS membrane by the introduction of rigid organophilic groups has potential in the pervaporation field for the improvement of separation behavior (Luo et al., 2008). Polyphenylmethylsiloxane–cellulose acetate (PPMS–CA) and polydimethylsiloxane–cellulose acetate (PDMS-CA) membranes are used for concentration of volatile organic compounds such as methanol, ethanol and acetone from aqueous solutions by pervaporation. The hydrophobicity of PPMS membranes is stronger than that of PDMS. In the study of Song and Lee (2005), the tube type alumina substrate was modified with silane coupling agent (perfluoroalkylsilane) to obtain hydrophobic membrane. The surface modified membrane showed much higher flux but lower selectivity than nonporous PDMS membrane for the recovery of esters. However, the selectivity of the membrane was sufficiently high. Olsson et al. (2002) studied the effect of three different POMS membranes, a polymethyloctylsiloxane–polyetherimide (POMS-PEI) and two polymethyloctylsiloxane–polyacrylonitrile (POMS-PAN) membranes having different thicknesses for the separation of alcohols, esters and aldehydes. They concluded that the porous support layer could affect the selectivity considerably and membranes should be designed with a low degree of crosslinking to obtain better separation properties, especially for larger permeants. © Woodhead Publishing Limited, 2010
230 Separation, extraction and concentration processes
8.5 Recovery of aroma compounds by pervaporation and applications in the food and beverage industries One of the most promising applications of pervaporation is the recovery of aroma compounds from aqueous mixtures. Aroma is a complex mixture of hundreds of different volatile organic compounds present at very low concentrations, typically at ppm or ppb levels. Aroma compounds belong to various chemical groups such as alcohols, aldehydes and esters. Recovery of natural flavor and aroma compounds has received much more attention in the food, biotechnology and cosmetic industries because natural aroma compounds are preferred to chemically synthesized ones. In the food industry, flavor concentrates are widely used to compensate for the loss of aroma compounds during processing. Because aroma compounds are highly heat sensitive and loss of aroma compounds during heat treatments is inevitable, they should be separated from food systems and added back to the final product. This can be accomplished either by recovering the lost aromas from the loss stream or by stripping the aromas from the raw material stream before processing. This eventually increases the quality and acceptance of the final product. There have been a large number of studies in this area. Pervaporation has been successfully applied to the recovery of the aroma compounds of wine (Karlsson et al., 1995; Schafer and Crespo, 2007; Schafer et al., 1999) and several fruit juices, such as apple (Bengtsson et al., 1992; Börjesson et al., 1996; Olsson and Tragardh, 1999 and 2001), cashew apple (De Assis et al., 2007), kiwifruit (Cassano et al., 2006), strawberry (Isci et al., 2006), pineapple (Pereira et al., 2002 and 2005; Sampranpiboon et al., 2000), passion fruit (Pereira et al., 2002; 2005), banana (Sampranpiboon et al., 2000), orange (Aroujalian and Raisi, 2007; Shepherd et al., 2002), blueberry (Peng and Liu, 2003), billberries (Diban et al., 2008; Garcia et al., 2008), grape (Rajagopalan and Cheryan, 1995) and pomegranate (Raisi et al., 2008). In addition, pervaporation has been found to be suitable for concentration of tea (Kanani et al., 2003), cocoa (Kattenberg and Willemsen, 2001) and dairy aroma compounds (Baudot and Marin, 1996; Overington et al., 2008) from aqueous solutions and also for dealcoholization of alcoholic beverages (Takacs et al., 2007). It was successfully applied for the recovery of terpenes from waste waters in essential oil industry (Charbit et al., 1997) and for deodorization of waste streams (Souchon et al., 2002). It can also be used as an analytical separation technique for the continuous determination of volatiles (Amador-Hernandez and Castro, 2000). In most of the studies, model feed solutions have been used because of the complexity of conducting the pervaporation experiments with real multicomponent mixtures (Baudot and Marin, 1999; Bengtsson et al., 1992; Börjesson et al., 1996; Garcia et al., 2008; Isci et al., 2006; Kanani et al., 2003; Olsson and Tragardh, 1999; 2001; Peng and Liu, 2003; Pereira et al., 2002 and 2005; Raisi et al., 2008; Rajagopalan and Cheryan, 1995;
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Principles of pervaporation for the recovery of aroma compounds 231
Sampranpiboon et al., 2000). However, there are also some studies in which real food samples are used (Alvarez et al., 2000; Kanani et al., 2003; Pereira et al., 2002 and 2005; Raisi et al., 2008; Rajagopalan and Cheryan, 1995; She and Huang, 2006). Most of the research on aroma recovery by pervaporation has been conducted on a laboratory scale. Studies of scale-up of the pervaporation process for the separation of multi-component mixtures are few. Karlsson et al. (1998) developed a model for the scale-up of pervaporation units for aroma recovery in which the effect of concentration polarization was taken into account with a semi-empirical model. A novel process simulation of pervaporation was developed for multicomponent mixtures (Lipnizki et al., 2002a) and this study was continued with the integration and optimization of hydrophobic pervaporation for the recovery of natural aroma compounds in the food industry (Lipnizki et al., 2002b). Trifunovic et al. (2006) investigated module design aspects of pervaporation using a modified version of an existing pervaporation simulation tool for aroma recovery. By applying the simulation to four aroma compounds: two alcohols (n-butanol, n-hexanol) and two esters (isoamyl acetate and ethyl butyrate), the effect of major process and module design parameters on the performance of a single module has been investigated. In general, for the compounds studied, the permeate composition can be manipulated by changing the module geometry and operating conditions because the effect of these parameters is more significant on esters, which in general have high recovery rates, than for alcohols, which are only marginally affected. 8.5.1 Recovery of fruit juice aromas Apple juice is one of the important fruit juices that pervaporation studies have focused on. Börjesson et al. (1996) investigated the performance of six different pervaporation membranes (PDMS-1060, PDMS-1070 (polydimethyl siloxane + silicalite), PDMS–PT 1100 (polydimethylsiloxane), POMS–PEI (polyoctylmethyl siloxane- polyethermide), POMS–PVDF and PEBA) for the recovery of apple juice aromas. Multicomponent solution containing five esters (ethyl acetate, ethyl butanoate, ethyl- 2-methylbutanoate, isopentyl acetate and hexyl acetate), one aldehyde (trans-2-hexenal) and four alcohols (isobutanol, butanol, isopentanol and hexanol) representing typical aroma compounds of apple juice was used in this study. The best performance was obtained with the PDMS–PT 1100, POMS–PEI and POMS–PVDF membranes. The influence of permeate pressure on the recovery of apple juice aroma compounds by pervaporation had also been examined using the same model solution (Olsson and Tragardh, 2001). A mathematical model, which includes the effect of concentration polarization on the feed side of the membrane and its dependence on permeate pressure, the variation in the separation properties of the membrane with permeate pressure and the properties of the individual permeants, was developed to predict the influence of permeate pressure on separation factor in pervaporation. © Woodhead Publishing Limited, 2010
232 Separation, extraction and concentration processes Moreover, the effect of feed flow velocity on recovery of apple juice aroma compounds by pervaporation was studied (Olsson and Tragardh, 1999). Increasing the feed flow velocity did not increase the recovery of the alcohols significantly because the dominating resistance to mass transfer was in the membrane for the alcohols. However, the recovery of esters and aldehydes was improved by increasing feed flow velocity. The dominating resistance to mass transfer was in the liquid feed boundary layer for the esters studied except ethyl acetate. For the aldehyde and ethyl acetate these two resistances were of the same order. Bengtsson et al. (1992) concentrated the twelve selected flavor compounds in apple juice by pervaporation. It was observed that alcohol had the lowest enrichment factors (2–22), the aldehydes had the medium values (16–67) and the esters had the highest enrichment factors (up to 100). The order of enrichment factors of aroma compounds of various chemical groups was the same in the study of Börjesson et al. (1996). De Assis et al. (2007) found that pervaporation can be used for the concentration of cashew apple juice aroma from cashew pulp. The chromatograms showed an increase in the peak area and also in the number of components in the permeate samples compared with the chromatogram of the cashew juice. Recovery of key flavor components from real flavor concentrates (apple essence, orange aroma and black tea distillate) was achieved by pervaporation in the study of She and Hwang (2006). In this study, both continuous and batch operations were carried out. Generally, the acetates and aldehydes exhibited higher enrichment factors than the alcohols. Significant flavor loss was observed in the pervaporation recovery process. PDMS membrane had higher permeation and loss rates than POMS membrane in batch pervaporation of dilute apple essence. Therefore, it was concluded that POMS membrane with a long operating time should be preferred if highly concentrated products and minimized flavor loss are the major concerns. However, PDMS membrane was found to be better if the time is limited. Rajagopalan and Cheryan (1995) studied pervaporation of model flavor compound of grapes (methyl anthranilate) and of commercial grape essence with several membranes: PDMS–PC (polydimethoxylsiloxane–polycarbonate), PEBA, PDMS-1070). PEBA membrane generally gave higher flux and better selectivity. Flux and selectivity decreased linearly with increase in downstream pressure, but increased with temperature. This relationship was not satisfied when the feed concentration was above 50 ppm where concentration polarization became serious. The presence of ethanol in the feed solution lowered separation factors but increased total flux. Experiments with commercial grape essence confirmed the excellent potential of pervaporation for the production of highly enriched flavors. Pervaporation was found to be a promising technique for the recovery of strawberry aroma compounds (Isci et al., 2006). The experiments were performed using binary aqueous solutions of methyl butyrate (MTB) and
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Principles of pervaporation for the recovery of aroma compounds 233
ethyl butyrate (ETB), ternary mixture of MTB, ETB and water and aqueous model strawberry solution containing six aroma compounds (MTB, ETB, butyl butyrate (BTB), methyl caproate (MTC), ethyl caproate (ETC) and linalool). An increase in both flux and selectivity was observed with an increase in feed temperature or decrease in downstream pressure in binary aqueous solution of MTB (Figs 8.3 and 8.4). Selectivity decreased with increase in feed concentration. Enrichment factors of MTB and ETB in binary and ternary solutions were not very different. However, MTB flux was retarded by half by the presence of ETB in ternary solution. In multicomponent mixtures of aqueous strawberry aroma solution, the selectivities of MTB and ETB were adversely affected by the presence of other compounds, owing to coupling effects. It is shown that flux coupling takes place when a permeant of low diffusivity is dragged through the membrane polymer by a permeant of higher diffusivity resulting in higher fluxes of the slower permeant than expected or the reverse. In the study of Sampranpiboon et al. (2000), pervaporation separation was used to recover aroma compounds from ethyl butanoate and ethyl hexanoate mixtures, which are the most important aroma components of pineapple and banana juice, by using POMS and PDMS membranes. The effect of operating conditions on the separation performance was investigated and in general the POMS membrane was found to be more selective to aroma compounds than PDMS membranes. Moreover, a strong interaction was observed between
Total mass flux (kg m–2 h–1)
0.3
0.2
0.1
0 20
30
40 Temperature (°C)
50
60
Fig. 8.3 Effect of feed temperature on total mass flux at different downstream pressures for 100 ppm methyl butyrate solution: ( ) 4 mbar, ( ) 8 mbar (reprinted from Isci et al., 2006, with permission from Elsevier).
© Woodhead Publishing Limited, 2010
234 Separation, extraction and concentration processes 120
100
Selectivity
80
60
40
20
0 20
30
40 Temperature (°C)
50
60
Fig. 8.4 Effect of feed temperature on selectivity of 100 ppm methyl butyrate solution at different downstream pressures: ( ) 4 mbar, ( ) 8 mbar (reprinted from Isci et al., 2006, with permission from Elsevier).
the two permeating components and permeation of one aroma compound was affected by the presence of the other aroma compound. Pineapple and passion fruit juices were chosen in the study of Pereira et al. (2002) and (2005). Performances of various membrane materials were evaluated with binary and quaternary synthetic aqueous solutions of typical tropical fruit aroma compounds such as ethyl acetate, ethyl butanoate, ethyl hexanoate and 1-octen-3-ol (Pereira et al., 2005). Tests were also carried out using single strength and clarified pineapple juices. Composite flat membranes having selective layers ethylene propylene diene terpolymer (EPDM) or ethylene vinyl acetate copolymer (EVA) and composite hollow fiber membrane having selective layer EPDM were used for comparison. It was concluded that choosing a very selective polymer is advantageous when the organic solute concentration is reduced in the feed. However, more permeable membranes are preferred if the feed concentration is high enough to induce phase separation in the permeate after its condensation. Composite EPDM hollow fiber membrane showed the best performance for pervaporation of synthetic and single-strength pineapple juices. Peng and Liu (2003), evaluated the separation factor of six aroma compounds (1-hexanol, 1-heptanol, trans-2-hexenal, ethyl acetate, linalool and d-limonene) representing typical flavoring ingredients from blueberry juice with a PDMS membrane. The results showed that the separation factor was in the range of 70 to 430, depending on molecule size and polarity of the compounds. No compounds, except 1-heptanol, showed any significant coupling effect in the mixture system. The temperature dependency of component fluxes was © Woodhead Publishing Limited, 2010
Principles of pervaporation for the recovery of aroma compounds 235
expressed by the Arrhenius equation and it was found that the temperature dependency of water flux was the highest whereas that of d-limonene was the lowest. Garcia et al. (2008) studied the concentration of trans-hex-2-en-1-ol, which is one of the major impact aroma compounds of bilberries, from aqueous ethanol using commercial PDMS capillary membranes in pervaporation. The flux of the aroma compound through the membrane was not affected from the variation of the feed flow rate. Therefore, it was concluded that the main mass transfer resistance to trans-hex-2-en-1-ol flux is located in the membrane. The temperature dependency of water, ethanol and trans-hex-2en-1-ol fluxes was expressed by the Arrhenius equation. From the magnitude of calculated activation energies, it was understood that water and ethanol were more influenced by variations in operation temperature than that of the aroma compound. Diban et al. (2008) adapted a previously developed mathematical model to predict the behavior of the separation of a mixture of seven volatile compounds that are characteristic of the aroma of bilberry juice from aqueous solution by pervaporation. It was observed that enrichment factors increased with membrane thickness until an asymptotic value was reached, however, partial fluxes decreased. Shepherd et al. (2002) studied the use of PDMS hollow fibers in orange juice aroma recovery. They used orange juice by product as feed and binary synthetic solutions. The well-spaced module was compared with a hollow fiber not containing spacers and with the transverse flow module with respect to their effect on experimental enrichment factors and mass transfer coefficient values and it was found to be feasible for aroma recovery. Aroujalian and Raisi (2007) also studied the pervaporation of orange juice aroma components. The effects of feed flow rate, feed temperature and permeate pressure on the performance of pervaporation was investigated. Feed flow rate had no significant effect on performance of the process. However, higher flux and enrichment factors were obtained as the feed temperature increased, whereas the increase in permeate pressure resulted in lower flux values. Enrichment factors of ethyl acetate, ethyl butyrate and hexanal increased but that of limonene, linalool and a-terpineol decreased as permeate pressure increased. Aroma compounds of pomegranate juice in multicomponent model solution and real fruit juice have been successfully concentrated by pervaporation (Raisi et al., 2008). The model aroma solution and real pomegranate juice showed similar behavior. The POMS membranes produced higher aroma enrichment factor but lower flux the PDMS membranes. Feed flow rate had no significant effect an either total flux on enrichment factor of aroma compounds, whereas on increase in feed temperature resulted in higher flux and enrichment factor. The effect of permeate pressure on pervaporation performance was related to the properties of the aroma compounds. Alvarez et al. (2000) investigated a combination of membrane processes
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236 Separation, extraction and concentration processes (enzymatic membrane reactor, reverse osmosis and pervaporation) for producing clarified juice and apple juice aroma concentrate. Cassano et al. (2006) performed a study to design an integrated membrane process for the production of concentrated kiwi juice and its aroma recovery. Pervaporation was carried out at different stages in the process. They suggested the use of pervaporation for the removal and enrichment of aroma compounds directly from fresh juice before any concentration process. 8.5.2 Recovery of wine aromas In literature, there are some studies on recovery of wine aromas by pervaporation (Karlsson et al., 1995; Schafer and Crespo, 2007; Schafer et al., 1999). Karlsson et al. (1995) carried out pervaporation with Muscat wine in which eight aroma compounds were identified. In this study, the increase in fluxes with feed temperature was very well described by the Arrhenius equation. The effects of temperature on permeation varied for different aroma compounds. Recovery of aroma compounds from a wine-must fermentation by organophilic pervaporation was studied by Schafer et al. (1999). Pervaporation was performed under fermentation conditions. It was shown that organophilic pervaporation can be highly suitable for the continuous recovery of very complex and delicate aromatic profiles produced during microbial fermentation. Schafer and Crespo (2007) determined the degree of concentration polarization in two different flow-cell configurations during recovery of winemust aroma compounds by pervaporation. It was found that concentration polarization could not be overcome even under turbulent feed flow conditions for compounds having high sorption coefficient for the respective membrane polymer. Dealcoholization of alcoholic beverages including beer and wine is another application of pervaporation. The idea behind this process is the separation of ethanol through hydrophobic membranes much more readily than water. There are many publications in this area (Brüschke, 1990; Escoudier et al., 1988; Kimmerle and Gudernatsch, 1991; Takacs et al., 2007). A problem encountered with this application is that the aroma compounds in the wine or beer are generally much more hydrophobic than ethanol and therefore permeate through the membrane even more readily than ethanol. Vankelecom et al. (1997) used a mixture of eight aromatic compounds in water, one of which was ethanol, in their study. Sorption of aromas in PDMS could be explained by the Hildebrandt solubility parameters. Sorption decreased when zeolites were added to the polymer owing to the cross-linking action of the zeolite. Silicalite-filled PDMS membrane was found to be the best for selective removal of ethanol from the aroma mixture. Therefore, high silicalite loading were recommended for preparation of alcohol-free beverages or removal of ethanol from biofermentations.
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Principles of pervaporation for the recovery of aroma compounds 237
Takacs et al. (2007) studied the reduction of alcohol content of wine by the pervaporation process using organophilic flat composite membrane, PDMS. Alcohol concentrate was obtained in the permeate stream. At higher temperatures, the flux of the permeate was higher and the membrane surface demand was smaller but the separation efficiency was lower. Lower pervaporation temperatures were more favorable to avoid aroma loss. In another study on the reduction of alcohol content of wine by pervaporation, the flux of the permeate increased but the separation efficiency and the separation ability of the membrane decreased as temperature increased (Takacs et al., 2007). In pervaporation of alcoholic beverages, complicated interactions between ethanol and aroma compounds often exist and affect their pervaporation performance. Tan et al. (2005) studied a series of model solutions containing ethanol and six typical aroma compounds in alcoholic beverages. The results showed that the solubility thermodynamics with the feed solution activity coefficient had a dominant effect on the permeability rather than diffusion factor. The presence of aroma compounds decreased the permeability coefficient and the separation factor of ethanol compared with those in ethanol–water binary solutions whereas its effect on fluxes was less. The effect of ethanol feed concentration on mass transfer of each compound was much related to the solubility properties of the compound in ethanol and water. It was observed that little interaction exists between aroma compounds in the membrane during the pervaporation process for dilute solution having a concentration below 1000 ppm. 8.5.3 Recovery of tea, cocoa and coffee aromas The possibility of using pervaporation to recover the tea aroma compounds from tea extract was studied by Kanani et al. (2003). Two different membrane types (POMS and PDMS) were used to recover eight aroma compounds that make a significant contribution to tea aroma. Pervaporation was performed with binary aqueous solutions of aroma compounds, with multicomponent mixture and with an actual tea extract. The actual tea aroma extract showed quite different behavior from the model aroma mixture. The results indicated that pervaporation is an attractive technology for the recovery of tea aroma compounds from tea aroma condensates. However, a wide range of selectivities for different aroma compounds results in alterations in the aroma profile of the permeate. Pilot plant tests for obtaining aroma extracts from cocoa using pervaporation were reported by Kattenberg and Willemsen (2001). The constructed test unit could process up to 500 L of feed, using hollow fiber membranes with a PDMS skin and two-stage condensation for permeate recovery. Application of pervaporation for recovering aroma concentrate from a caffeine- or theobromine-containing food such as coffee or tea and, in particular, cocoa was investigated by Kattenberg et al. (2002). In this study, first an
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238 Separation, extraction and concentration processes aqueous extract was obtained. After incubating the food material at a suitable temperature and for a suitable time, the food extract was pervaporated using a hydrophobic pervaporation membrane to obtain the aroma concentrate. 8.5.4 Recovery of dairy aromas It is also possible to concentrate dairy flavor compounds from aqueous solutions using pervaporation. Baudot and Marin (1996) concentrated two dairy aroma compounds (methylthiobutanoate, diacetyl) in model aqueous solutions by pervaporation through two different membranes (silicalitefilled silicone composite membrane (PDMS 1070) and PEBA homogeneous membrane. Tested membranes showed good selectivity for the extraction of methylthiobutanoate (hydrophobic cheese aroma) at high dilution rate. However, these membranes were less selective for the recovery of diacetyl (2,3-butanedione, hydrophilic butter aroma). For this component, two-stage condensation coupled with pervaporation improved the selectivity of the whole process significantly. Overington et al. (2008) worked with the model feed solution containing nine flavor compounds contributing the flavor of cheese and other dairy products in pervaporation. They studied the effect of membrane type (two types of PDMS and a POMS), feed temperature and permeate pressure. Total flux increased with increase in temperature but decreased with increase in permeate pressure. PDMS membranes gave higher total flux than POMS membrane but the POMS membrane gave higher enrichment factors for the major compounds in the permeate. Esters and ketones passed through the membrane more readily than acids. Diffusion was the controlling mechanism for esters and ketones. Enrichment factors decreased with increasing molecular weight within esters and ketones. However, the effect of molecular weight was more complex for acids and depended on the relative importance of sorption and diffusion mechanisms in the membrane. 8.5.5 Pervaporation as an analytical separation technique Analytical separation techniques play an important role in analytical chemistry. Analytical pervaporation has been successfully used for the continuous determination of volatile analytes or volatile reaction products from complex liquid, semi-solid and solid samples in food and beverage industries (Amador-Hernandez and Castro, 2000). Schafer et al. (2006) investigated the potential of pervaporation combined with an electronic nose based on metal oxide sensors for analyzing wine model solutions. It was shown that target solutes present at minor concentrations can be detected even in the presence of bulk interferents such as ethanol by combining pervaporation with an artificial olfactory system based on metal oxide sensors.
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Principles of pervaporation for the recovery of aroma compounds 239
8.5.6 Other applications of pervaporation for the recovery of aromas Pervaporation can also be used for the extraction of aroma compounds from natural matrixes as an alternative to steam distillation and solvent extraction. Enzymatic pretreatment for degradation of cell membrane is necessary in these processes. Figoli et al. (2006) studied the recovery of bergamot peel oil by pervaporation. There are some studies that combine pervaporation with bioprocessing. Bluemke and Schrader (2001) carried out an integrated bioprocess for production and recovery of the synthesized aroma compounds by interlinking a pervaporation membrane (POMS, PEBA) module with a bioreactor. In this study, they used the fungus Ceratocystis moniliformis which produce ethyl acetate, propyl acetate, isobutyl acetate, isoamyl acetate, citronellol and geraniol. In situ product removal using pervaporation leads to decreased product concentrations in the bioreactor and increased microbial growth rates. Total yield of aroma compounds produced was higher in the integrated bioprocess than with batch cultivation. Böddeker et al. (1997) showed that natural vanillin formed by bioconversion of suitable precursors can be recovered directly from the acidified culture broth by pervaporation at elevated temperature. Large volumes of waste water are obtained in most food processes. Since this liquid is generally released into the environment, it represents both pollution and an economical loss because of the cost of the valuable organic molecules involved. Pervaporation can be used to remove these valuable aroma compounds from waste water. Pierre et al. (2001) carried out non-dispersive solvent extraction of three sulfur aroma compounds, dimethyldisulfide, dimethyltrisulfide and S-methyl thiobutanoate, from very dilute aqueous solutions representing real effluent. Mass transfer fluxes obtained experimentally by membrane-based solvent extraction were greater for the three aroma compounds than those obtained by pervaporation. Souchon et al. (2002) used pervaporation process for deodorization of cauliflower blanching effluent in order to reduce its volatile organic compounds content and to recover the valuable food flavoring fraction. Pervaporation has been performed on three sulfur compounds of the cauliflower odor. Pervaporation was successfully applied for the recovery of linalool and eucalyptol which are the most frequent terpenes in wastewaters from the essential oil industry (Charbit et al., 1997).
8.6 Sources of further information and future trends Pervaporation theory and principles, separation characteristics, thermodynamics and polymer materials for membranes are covered in the book edited by Huang (1991), in which the chapter by Böddeker and Bengtson (1991) deals with the separation of organics from water by pervaporation, and that
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240 Separation, extraction and concentration processes by Koops and Smolders (1991) gives information about membranes used in pervaporation. Strathmann and McDonogh (1993) studied mass transfer in pervaporation, characterization of pervaporation membranes and use of pervaporation in biotechnology. Brüschke (2006) studied the transport resistances and principles of pervaporation, membranes and the application of pervaporation in the chemical industry. Mass transport phenomena during the recovery of volatile compounds by pervaporation were investigated by Schafer and Crespo (2002). Pereira et al. (2006) reviewed the applications of pervaporation for the recovery of volatile aroma compounds from fruit juices. The studies on scale-up of the pervaporation process are limited and mostly based on the separation of model solutions. Therefore, pervaporation studies using real food systems should be increased in future. In addition, further study is required on scale-up. Studying pervaporation in large-scale systems provides insight into the industrial application of this method. Future trends are towards the combination of pervaporation with other processes in the food process line.
8.7 References Alvarez S, Riera F A, Alvarez R, Coca J, Cuperus F P, Bouwer S T, Boswinkel C, van Gemert RW, Veldsink JW, Giorno L, Donato L, Todisco S, Drioli E, Olsson J, Tragardh G, Gaeta S N and Panyor L (2000), ‘A new integrated membrane process for producing clarified apple juice and apple juice aroma concentrate’, J Food Eng, 46, 109–125. Amador-Hernandez J and de Castro MDL (2000), ‘Pervaporation: a useful tool in food analysis’, Food Chem, 68, 387–394. Aroujalian A and Raisi A (2007), ‘Recovery of volatile aroma components from orange juice by pervaporation’, J Membr Sci, 303, 154–161. Baudot A and Marin M (1996), ‘Dairy aroma compounds recovery by pervaporation’, J Membr Sci, 120, 207–220. Bengtsson E, Tragardh G and Hallstrom B (1992), ‘Concentration of apple juice aroma from evaporator condensate using pervaporation’, Lebensm Wiss Technol, 25, 29–34. Bluemke W and Schrader J (2001), ‘Integrated bioprocess for enhanced production of natural flavors and fragrances by Ceratocystis moniliformis’, Biomol Eng, 17, 137–142. Borjesson J, Karlsson H O E and Tragardh G (1996), ‘Pervaporation of a model apple juice aroma solution: comparison of membrane performance’, J Membr Sci, 119, 229–239. Böddeker K and Bengtson G (1991), ‘Selective pervaporation of organic from water’, in R. Y. M. Huang, Pervaporation membrane separation processes, Elsevier, Amsterdam, The Netherlands, 437–460. Böddeker K W, Gatfield I L, Jahnig J, Schorm C (1997), ‘Pervaporation at the vapor pressure limit: Vanillin’, J Membr Sci, 137, 155–158. Brüschke H E A (1990), ‘Removal of ethanol from aqueous streams by pervaporation’, Desalination, 77, 323–329. Brüschke H E A (2006), ‘State-of-art of pervaporation process in the chemical industry’, in S. P. Nunes and K. V. Peinemann, membrane technology in the chemical industry, Wiley-VCH, USA, 151–202.
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Principles of pervaporation for the recovery of aroma compounds 241 Cassano A, Figoli A, Tagarelli A, Sindona G, Drioli E (2006), ‘Integrated membrane process for the production of highly nutritional kiwifruit juice’, Desalination, 189, 21–30. Charbit G, Charbit F and Molina C (1997), ‘Study of mass transfer limitations in the deterpenation of waste waters by pervaporation’, J Chem Eng Jpn, 30, 382–387. De Assis A V R, Bizzo H R, da Matta V M, Cabral L M C (2007), ‘Recovery of aroma compounds of cashew apple fruit (Anacardium accidentale L.) by pervaporation’, Ciencia E Technologia De Alimentos, 27, 349–354. Diban N, Urtiaga A and Ortiz I (2008), ‘Recovery of key components of bilberry aroma using a commercial pervaporation membrane’, Desalination, 224, 34–39. Escoudier J L, Le Bouar M, Moutounet M, Jouret C and Barillere J M (1988), ‘Application and evaluation of pervaporation for production of low alcohol wines’, in R. Bakish, Proceedings of the Third International Conference on Pervaporation Processes in the Chemical Industry, Bakish Materials Corporation, Englewood Cliffs, NJ, USA, 387–397. Figoli A, Tagarelli A, Mecchia A, Trotta A, Cavaliere B, Lavecchia R, Sindona G and Drioli E (2006), ‘Enzyme-assisted pervaporative recovery of concentrated bergamot peel oils’, Desalination, 199, 111–112. Garcia V, Diban N, Gorri D, Keiski R, Urtiaga A and Ortiz I (2008), ‘Separation and concentration of bilberry impact aroma compound from dilute model solution by pervaporation’, J Chem Technol Biotechnol, 83, 973–982. Huang R Y M (1991), Pervaporation membrane separation processes, Elsevier Science Publishers, Amsterdam, The Netherlands. Isci A, Sahin S and Sumnu G (2006), ‘Recovery of strawberry aroma compounds by pervaporation’, J Food Eng, 75, 36–42. Kanani D M, Nikhade B P, Balakrishnan P, Singh G and Pangarkar V G (2003), ‘Recovery of Valuable Tea Aroma Components by Pervaporation’, Ind Eng Chem Res, 42, 6924–6932. Karlsson H O E and Tragardh G (1993), ‘Aroma compound recovery with pervaporationfeed flow effects’, J Membr Sci, 81, 163–171. Karlsson H O E, Loureiro S and Tragardh G (1995), ‘Aroma compound recovery with pervaporation-Temperature effects during pervaporation of a muscat wine’, J Food Eng, 26, 177–191. Karlsson H O E, Tragardh G and Olsson J (1998), ‘The performance of pervaporative aroma recovery units: Process simulations’, Sep Sci Technol, 33, 1629–1652. Kattenberg H R and Willemsen J H A (2001), ‘Aroma extracts from cocoa’, Manuf Confect, 82, 73. Kattenberg H R, Willemsen J H A, Starmans D A J, Hoving H D and Winters MGM (2002), ‘Method for recovering aroma concentrate from caffeine- or theobrominecomprising food base material’, US Patent Office Journal, No. 1139792. Kimmerle K and Gudernatsch W (1991), ‘Pilot dealcoholization of beer by pervaporation’, in R. Bakish, Proceedings of the Fifth International Conference on Pervaporation Processes in the Chemical Industry, Bakish Materials Corporation, Englewood Cliffs, NJ, USA, pp. 291–307. Kober P A (1917), ‘Pervaporation, perstillation, and percrystallisation’, J Am Chem Soc, 39, 944–948. Koops G H and Smolders C A (1991), ‘Estimation and evaluation of polymeric materials for pervaporation membranes’, in R. Y. M. Huang, Pervaporation membrane separation processes, Elsevier, Amsterdam, The Netherlands, 253–278. Lipnizki F, Hausmanns S, Ten P, Field R W and Laufenberg G (1999), ‘Organophilic pervaporation: prospects and performance’, Chem Eng J, 73, 113–129. Lipnizki F, Olsson J and Tragardh G (2002a), ‘Scale-up of pervaporation for the recovery of natural aroma compounds in the food industry. Part 1: Simulation and performance’, J Food Eng, 54, 183–195.
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242 Separation, extraction and concentration processes Lipnizki F, Olsson J and Tragardh G (2002b), ‘Scale-up of pervaporation for the recovery of natural aroma compounds in the food industry. Part 2: Integration and optimization’, J Food Eng, 54, 197–205. Luo Y, Tan S J, Wang H, Wu F W, Liu X M, Li L, Zhang Z B (2008), ‘PPMS composite membranes for the concentration of organics from aqueous solutions by pervaporation’, Chem Eng J, 137, 496–502. Olsson J and Tragardh G (1999), ‘Influence of feed flow velocity on pervaporative aroma recovery from a model solution of apple juice aroma compounds’, J Food Eng, 39, 107–115. Olsson J and Tragardh G (2001), ‘Pervaporation of volatile organic compounds from water. I. Influence of permeate pressure on selectivity’, J Membr Sci, 187, 23–37. Olsson J, Tragardh G and Lipnizki F (2002), ‘The influence of permeant and membrane properties on mass transfer in pervaporation of volatile organic compounds from dilute organic solutions’, Sep Sci Technol, 37, 1199–1223. Overington A, Wong M, Harrison J and Ferreira L (2008), ‘Concentration of dairy flavour compounds using pervaporation’, Int Dairy J, 18, 835–848. Peng M and Liu S X (2003), ‘Recovery of aroma compounds from dilute model blueberry solution by pervaporation’, J Food Sci, 68, 2706–2710. Pereira C C, Ribeiro C P, Nobrega R, Borges C P (2006), ‘Pervaporative recovery of volatile aroma compounds from fruit juices’, J Membr Sci, 274, 1–23. Pereira C C, Rufino J M, Habert A C, Nobrega R, Cabral L M C and Borges C P (2002), ‘Membrane for processing tropical fruit juice’, Desalination, 148, 57–60. Pereira C C, Rufino J R M, Habert A C, Nobrega R, Cabral L M C and Borges C P (2005), ‘Aroma compounds recovery of tropical fruit juice by pervaporation: membrane material selection and process evaluation’, J Food Eng, 66, 77–87. Pierre F X, Souchon I, Marin M (2001), ‘Recovery of sulfur aroma compounds using membrane-based solvent extraction, J Membr Sci, 187, 239–253. Raisi A, Aroujalian A, Kaghazchi T (2008), ‘Multicomponent pervaporation process for volatile aroma compounds recovery from pomegranate juice’, J Membr Sci, 322, 339–348. Rajagopalan N and Cheryan M (1995), ‘Pervaporation of grape juice aroma’, J Membr Sci, 104, 243–250. Sampranpiboon P, Jiraratananon R, Uttapap D, Feng X and Huang R Y M (2000), ‘Separation of aroma compounds from aqueous solutions by pervaporation using polyoctylmethyl siloxane (POMS) and polydimethyl siloxane (PDMS) membranes’, J Membr Sci, 174, 55–65. Schafer T and Crespo J G (1997), ‘Recovery of aromas by pervaporation’, ESMST XIVth Annual Summer School, Membrane integration in clean processes. Lisbon Instituto: Superior Tecnico, IST. Schafer T and Crespo J G (2002), ‘Mass transport phenomena during the recovery of volatile compounds by pervaporation’, in J. Welti-Chanes, J. F. Velez-Ruiz and G.V. Barbosa-Canovas, Transport phenomena in food processing, CRC Press, Florida, USA, 247–263. Schafer T and Crespo J G (2007), ‘Study and optimization of the hydrodynamic upstream conditions during recovery of a complex aroma profile by pervaporation’, J Membr Sci, 301, 46–56. Schafer T, Bentson G, Pingel H, Boddeker K W and Crespo J P S G (1999), ‘Recovery of aroma compounds from a wine-must fermentation by ogranophilic pervaporation’, Biotechnol Bioeng, 62, 412–421. Schafer T, Serrano-Santos M B, Rocchi S and Fuoco R (2006), ‘Pervaporation membrane separation process for enhancing the selectivity of an artificial olfactory system (“electronic nose”)’, Anal Bioanal Chem, 384, 860–866.
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Principles of pervaporation for the recovery of aroma compounds 243 She M and Hwang S-T (2006), ‘Recovery of key components from real flavor concentrates by pervaporation’, J Membr Sci, 279, 86–93. Shepherd A, Habert A C and Borges C P (2002) ‘Hollow fiber modules for orange juice aroma recovery using pervaporation’, Desalination, 148, 111–114. Song K H and Lee K R (2005), ‘Pervaporation of flavors with hydrophobic membrane’, Korean J Chem Eng, 22, 735–739. Souchon I, Pierre F X, Athes-Dutour V and Marin M (2002), ‘Pervaporation as a deodorization process applied to food industry effluents: recovery and valorisation of aroma compounds from cauliflower blanching water’, Desalination, 148, 79–85. Strathmann H and McDonogh R M (1993), ‘Use of pervaporation in biotechnology’, in J. A. Howell, V. Sanchez and R. W. Field, Membranes in bioprocessing: theory and applications, Chapman and Hall, Cambridge, UK, 293–329. Takacs L, Vatai G, Korany K (2007), ‘Production of alcohol free wine by pervaporation’, J Food Eng, 78, 118–125. Tan S J, Li L, Xiao Z Y, Wu Y T and Zhang Z B (2005), ‘Pervaporation of alcoholic beverages – the coupling effects between ethanol and aroma compounds’, J Membr Sci, 264, 129–136. Trifunovic O, Lipnizki F and Tragardh G (2006), ‘The influence of process parameters on aroma recovery by hydrophobic pervaporation’, Desalination, 189, 1–12. Vankelecom I F J, Beukelaer S D and Uytterhoeven J B (1997), ‘Sorption and pervaporation of aroma compounds using zeolite-filled PDMS membranes’, J Phys Chem B, 101, 5186–5190.
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244 Separation, extraction and concentration processes
9 Advances in membrane-based concentration in the food and beverage industries: direct osmosis and membrane contactors E. Drioli and A. Cassano, Institute on Membrane Technology, ITM-CNR, Italy Abstract: The process fundamentals and the role of operating conditions involved in direct osmosis (DO) and membrane contactors, such as osmotic distillation (OD) and membrane distillation (MD) are reviewed. Their main applications in concentration of liquid food and limitations compared with those of conventional de-watering processes – such as thermal evaporation, cryoconcentration and reverse osmosis – are discussed. The advantage lies in allowing very high concentrations (up to 65 °Brix) to be reached under atmospheric pressure and low temperature, thus avoiding thermal and mechanical damage. Key words: membrane contactors, direct osmosis, osmotic distillation, membrane distillation, food industry.
9.1 Introduction Epidemiological studies report that high consumption of fruit and vegetables is associated with a reduced risk of free radical related oxidative damage and diseases, such as various types of cancer, and cardiovascular or neurological diseases (Collins and Harrington, 2002). The world market of these products has remarkably increased in recent years and consumers have addressed their interest towards products with a natural fresh taste, that are high quality and additive-free. In order to meet these demands, the food industry has focused on the development of processed items with increased shelf-life able to retain as much as possible the essence of the fresh fruit, as well as colour, aroma, nutritional value and structural characteristics.
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Advances in membrane-based concentration 245 In the traditional methods of processing liquid foods, the concentration operation is particularly important for the production of intermediates that can be easily handled, packaged, stored and shipped before their reutilization for a final transformation (Luh et al., 1986) and also as a step which precedes total dehydration. Moreover, the removal of water reduces the growth of micro-organisms, thus increasing the shelf-life of the liquid foods; fruit juice concentrates, because of their low water activity, have a higher stability than single-strength juices. This chapter briefly describes the existing methods for the concentration of liquid foods (such as thermal evaporation, cryoconcentration and membrane concentration) and then introduces novel membrane processes, including direct osmosis (DO), membrane distillation (MD) and osmotic distillation (OD), which offer improved performances in terms of product quality, reduced energy consumption and environmental impact, and high effectiveness.
9.2 Conventional technologies in the food and beverage industries 9.2.1 Thermal evaporation In thermal evaporation, water is removed from liquid foods as vapour. Falling film evaporators and centrifugal evaporators are most commonly used for this purpose. It is known that the heat required to perform the evaporation results in some ‘cooked’ notes recognized as off-flavours, loss of most volatile aroma compounds, and colour degradation with a consequent remarkable qualitative decline (Maccarone et al., 1996). Partial recovery of aroma compounds may be achieved by cold-water condensation and rectification in a still located on the evaporator, resulting in concentrates that are acceptable to consumers, but still far from fresh products. An additional drawback in the use of thermal evaporation is the high energy demand, despite the use of energy-saving systems. The production of superior quality concentrates, especially fruit juices, is a very important goal and considerable R&D efforts have been devoted to the development of nonthermal concentration techniques. These methods include: freeze concentration systems (cryoconcentration) and membrane processes. 9.2.2 Cryoconcentration Freeze concentration has long been considered a feasible prospect for the selective removal of water from liquid foods. In this instance, water is removed as ice rather than as vapour. The product is cooled below its freezing point allowing water to form crystals which are then removed from the concentrate by centrifugation. This method guarantees an excellent organoleptic quality: the operating temperatures are sufficiently low to avoid chemical and
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246 Separation, extraction and concentration processes biochemical reactions and volatile compounds are completely retained. Consequently, the concentrated product is characterized by higher quality than that processed by thermal evaporation (Aider and de Halleux, 2008). However, there are some problems in this process which make it unpractical. The main drawback is that the achievable concentration (40–45 °Brix) is lower than the values obtained by evaporation (60–65 °Brix). Furthermore, fine ice crystals produced by rapid freezing cannot be separated from the residual liquid and the process is not suitable for the treatment of highly pulped liquids. Remarkable energy consumption, high equipment investments and operational costs are additional drawbacks (Jariel et al., 1996). 9.2.3 Pressure-driven membrane processes Pressure-driven membrane processes such as microfiltration (MF), ultrafiltration (UF), nanofiltration (NF) and reverse osmosis (RO) are well established systems in the food industry for concentration, fractionation and purification of liquid foods and for wastewater treatment (Table 9.1). These processes offer significant advantages over traditional food processing technologies in terms of low operation temperatures, no special chemicals required, possibility of automation, and no phase-change involved. In addition, they are simple in concept and operation and characterized by low energy consumption. MF membranes have a symmetric microporous structure with a pore size in the range 0.1–10 mm. They are essentially used for the separation of fine particles, micro-organisms and emulsion droplets from fluids. Particles are rejected mainly by means of sieving mechanisms although the separation is affected also by interactions between the membrane surface and the treated solution. The hydrostatic pressure difference used as a driving force is in the range from 0.5 to 4 bar. Cell harvesting, clarification of fruit juices, must and wine, wastewater treatment, bacteria and particulate turbidity reduction, separation of whey proteins and separation of oil–water emulsions are typical consolidated applications (Cheryan, 1998). UF membranes have a pore size of 1–100 nm and are capable of retaining species of molecular weight 300–500 000 daltons. Typical rejected species include lipids, proteins, polysaccharides, colloids and suspended solids whereas solutes with small molecules such as vitamins, salts and sugars, flow through the membrane together with water. Operating pressures are 2–10 bar (Mulder, 1998). In the food and dairy industry, UF is commonly applied to clarify, concentrate, decolorize and fractionate juices, sugar solutions, proteins and dairy and grain milling products. The membranes used in NF are characterized by a charged surface and equivalent pore diameters in the range 1–3 nm. They are mainly used for the separation of multivalent ions and uncharged organic molecules having molecular weights greater than 1 kDa. Operating
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Advances in membrane-based concentration 247 Table 9.1 Commercial application of pressure-driven membrane processes in food processing Unit operation Industrial sector
Application
Microfiltration Fruit and vegetables
Fruit and vegetable juice clarification
Beverages Dairy
Ultrafiltration
Fruit and vegetables Beverages Vegetables Dairy
Fats and oils Nanofiltration
Fruit and vegetables Dairy Beverages
Reverse osmosis
Cold sterilization of beer Clarification of wine, beer and vinegar Bacterial purification of milk and whey Whey defatting and high-quality protein recovery Clarification of fruit juice Clarification of wine Recovery of potato starch and proteins Fractionation/concentration of whey in manufacture of whey protein concentrate Fractionation/concentration of milk in manufacture of processed and natural cheese Recovery of oilseed protein and oil from processing water Fruit and vegetable juice concentration Concentration and demineralization of milk Demineralization of whey Wastewater treatment
Fruit and vegetables
Fruit and vegetable juice concentration Sugar concentration and recovery from rinse water, treatment of water for reuse Recovery of flavours, fragrances, pectins and proteins Beverages Recovery of alcohol from wine Sugar and sweeteners Preconcentration of maple syrup Sugar concentration and recovery from rinse water, treatment of water for reuse Dairy Concentration of milk and whey All possible Treatment of boiler concentrate prior to recycle to boiler
pressures in NF are lower than those in RO therefore separation occurs at low energy consumption (21% less than RO). The concentration of must, the recovery of aromas from fruit juices, the demineralization of whey and the treatment of wastewater from beverage production are typical applications of NF in the food and dairy industry (Warczok et al., 2004). In RO, low molecular weight compounds, such as salts and sugars are separated from a solvent, usually water. The particle size range for application of RO is approximately 0.1–1 nm and a complete separation is achieved with solutes of molecular weight greater than 300 Da. Typical operating pressures are 10–100 bar. Food processing applications of RO include: concentration
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248 Separation, extraction and concentration processes of fruit and vegetable juices, concentration of milk and whey for cheese production, pretreatment of boiler water, recovery of sugars and acids from rinse waters, recovery of sweet potato stillage, recovery of flavours, pectins and proteins, removal of alcohol from wine, and water softening (Ho and Sirkar, 1992). The advantages of RO over traditional evaporation are the lower thermal damage to the product, increase in aroma retention, simple system design, lower energy consumption and lower equipment costs. However, RO cannot reach retentate concentrations higher than 25–30 °Brix with a single-stage system owing to the limitation of high osmotic pressure. Furthermore, fouling and concentration polarization are still important drawbacks of RO and other pressure-driven membrane processes in liquid food processing (Jiao et al., 2004).
9.3 Direct osmosis and applications in the food and beverage industries 9.3.1 Process fundamentals In DO, an osmotic pressure gradient across a semi-permeable membrane is established by using an osmotic agent solution in order to remove water from a feed solution. The osmotic agent should be characterized by high solubility in water (low water activity), non-toxicity, high superficial tension, inertia towards the flavour, odour and colour of the foodstuff, low volatility and viscosity, and non-permeability across the membrane. The most frequently employed constituents as osmotic agents are: sodium chloride, calcium chloride, glucose, sucrose, glycerol, cane molasses or corn syrup. Unlike RO the pressure difference across the membrane in DO is negligible and the flux depends on the difference in water activities. The hydraulic pressures required to pump the feed and the osmotic solution over the membrane surface are of about 2 bar (Milleville, 1990). Consequently, fouling phenomena are negligible and only concentration polarization occurs. Additional advantages are low energy consumption, simplicity, modularity, constant permeate flux in time, possibility to treat solutions with a high level of suspended and dissolved solids, and high achievable concentrations. The membranes used in DO are semipermeable membranes similar to those used in RO and NF processes. They are highly selective; however, the diffusion of a small amount of stripping solution cannot be completely avoided. 9.3.2 Effect of operating conditions on the direct osmosis (DO) flux Petrotos et al. (1998, 1999) studied the effect of several membranes and operating parameters on the performance of DO in the concentration of
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Advances in membrane-based concentration 249 tomato juice. NaCl, CaCl2, Ca(NO3)2, glucose, sucrose and polyethylene glycol (PEG 400) were used as stripping solutions. The initial concentration of tomato juice was in the range 4.3–11.7 °Brix. These studies suggested that salts, and especially NaCl, produced improved fluxes compared with carbohydrates and PEG 400, owing to their lower viscosity. An increasing of the stripping solution concentration also positively affected the flux (through an increasing of the driving force). On the contrary, an increasing of the feed concentration resulted in lower water flux owing to viscosity and osmotic pressure increase. The lower viscosity of the osmotic agents resulted in reduced polarization phenomena on the stripping solution side and increased fluxes. Increasing the temperature from 26 to 60 °C enhanced permeate flux by 64%, this phenomenon can be attributed to the reduction in the viscosity and to the increase in the diffusion coefficient of both feed and stripping solution. An increase in feed flow rate did not improve the permeate flux during the DO process. The pretreatment of the juice by MF or UF was also found to have a significant effect on the osmotic fluxes (Petrotos et al., 1998). In particular, the ultrafiltered juice gave a 39% increase in flux compared with the untreated juice. Finally, the DO flux is strongly affected by membrane thickness. In particular, a reduction of the membrane support layer determines an increase of the permeate flux (Beaudry and Lampi, 1990). 9.3.3 DO applications Osmotek Inc. (Corvallis, OR, USA) developed the first DO industrial approach as a pretreatment step for RO wastewater treatment. Several studies concerning the fruit juice concentration by DO are reported (Beaudry and Lampi, 1990; Herron et al., 1994; Milleville, 1990; Petrotos and Lazarides, 2001; Petrotos et al., 1998, 1999; Wrolstad et al., 1993). Beaudry and Lampi (1990) reported the performance of the Osmotek DO process in the concentration of fruit juices by using thin-film composite RO membranes with an overall thickness of 25–85 mm and a nominal molecular weight cut off (NMWCO) of 100 Da. These membranes rejected more than 99.9% of the total acidity and colour in orange and red raspberry juice. Herron et al. (1994) developed a simplified DO apparatus for several kinds of liquid food samples including orange and raspberry juice. Maximum osmotic fluxes of 5–6 L m–2 h–1 were obtained through 100-Da membranes by using fructose/glucose solutions as osmotic agent. The concentrated juice showed a better quality than that produced by a conventional vacuum evaporator. Stripping solutions diluted during juice concentration can be reconcentrated by using thermal evaporation: therefore, only the osmotic agent is exposed to high temperatures preserving the quality of concentrated juice. RO was also proposed as a method for the recovery of stripping solutions (Karode et al., 2000). Petrotos and Lazarides (2001) developed a new apparatus for the concentration of fruit juices with an osmotic cell of special configuration
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250 Separation, extraction and concentration processes to promote turbulence. The apparatus was equipped with RO flat-sheet membranes to concentrate tomato juice up to 16 °Brix at room temperature. Average fluxes of 4.5 kg m–2h–1 were obtained.
9.4 Membrane contactors and applications in the food and beverage industries Membrane contactors are devices in which a microporous membrane acts as a barrier between two phases permitting gas/liquid or liquid/liquid mass transfer of the components without dispersion of one phase into the other. Fluids to be contacted flow on the opposite side of the membrane and a fluid/fluid interface is located at the mouth of each membrane pore. Mass transfer occurs by diffusion across the interface as in conventional contacting equipment. Unlike conventional membrane operations, the driving force for separation is a concentration or a temperature gradient rather than a pressure gradient and a small pressure drop across the membrane is required to maintain the fluid/fluid interface at the mouth of the pore. Furthermore, the membrane imparts no selectivity to the separation (Gabelman and Hwang, 1999). Membrane contactors offer significant advantages over conventional dispersed-phase contactors such as the absence of emulsions, no density difference required between the fluids, easy scale-up, high interfacial area, no flooding at high flow rates. Within membrane contactors, membrane distillation (MD) and osmotic distillation (OD) seem to be a valid alternative for pressure-driven membrane processes in the concentration of liquid foods. In these processes, schematically represented in Fig. 9.1, the driving force for mass transfer is a vapour pressure difference across the membrane generated by either a temperature gradient (in MD) or a water activity difference (in OD). Their main advantages over traditional pressure-driven membrane processes are in terms of: low fouling, possibility of treatment of highly viscous solutions, high retention of species, and low energy consumption. Furthermore, these processes are not limited by high osmotic pressures and allow concentration levels to be attained that are similar to those obtained in thermal evaporation. 9.4.1 Osmotic distillation Process fundamentals Osmotic distillation is a recent membrane process also known as osmotic evaporation, osmotic concentration by membrane, membrane evaporation, isothermal membrane distillation, transmembrane distillation, thermopervaporation or gas membrane extraction. It can be carried out at room temperature and atmospheric pressure with minimal thermal and mechanical damage of the solutes. Consequently, it represents an attractive process for © Woodhead Publishing Limited, 2010
Advances in membrane-based concentration 251 Evaporation
Condensation
Tf > Tp Vapour
L/G interphase
Hydrophobic membrane
L/G interphase
Feed
Permeate (a)
Evaporation
Condensation
awf > aws Vapour
L/G interphase
Hydrophobic membrane
Feed
L/G interphase Stripping solution
(b)
Fig. 9.1 Schematic representation of (a) membrane distillation and (b) osmotic distillation. L/G, liquid–gas.
the concentration of solutions containing thermosensitive compounds, such as fruit juices and pharmaceuticals. In OD, a hydrophobic microporous membrane separates two liquid phases having different solute concentrations: a dilute solution on one side and a hypertonic salt solution on the opposite side. The hydrophobic nature of the membrane prevents penetration of the pores by aqueous solutions, creating air gaps within the membrane. The difference in solute concentration and, consequently, in water activity of both solutions, generates, at the vapour–liquid interface, a vapour pressure gradient across the membrane causing a vapour transfer across the pores from the high-vapour pressure phase to the low one (Alves and Coelhoso, 2002; Hogan et al., 1998; Lebfevre, 1988; Mengual et al., 1993). The concentration profile of the OD process is schematically presented in Fig. 9.2. The water transport through the membrane can be summarized in three steps: (1) evaporation of water at the dilute vapour–liquid interface; (2) diffusional vapour transport through the membrane pore; (3) condensation of water vapour at the membrane/brine interface (Schofield et al., 1987; Hogan et al., 1998). © Woodhead Publishing Limited, 2010
252 Separation, extraction and concentration processes High water vapour pressure Low water vapour pressure
Vapour flux
Feed dilutes aqueous solution
Csb Pwf >> Pwp Pore
Csm
Stripping solution (concentrated salt solution)
Cfm Cfb
Air Microporous hydrophobic membrane
Bulk Boundary layer
Bulk Boundary layer
Fig. 9.2 Concentration profile in osmotic distillation.
Because water transport involves condensation and evaporation phenomena, a temperature gradient through the membrane is generated, even if bulk temperatures of solutions separated by the membrane are equal. Consequently, a heat transfer tending to reduce the driving force for the water transport should be considered in addition to a mass transfer (Celere and Gostoli, 2002). The stripping solution after its dilution by water transferred from the feed stream can be reconcentrated by evaporation and reused in the OD operation. Therefore it must be thermally stable and preferably nontoxic, noncorrosive and available at low cost. A number of water-soluble salts such as NaCl, CaCl2, K2HPO4 are suitable. Salts displaying increase in solubility with temperature are preferred because they can be concentrated to high levels avoiding crystallization phenomena. Potassium salts of orthoand pyrophosphoric acid offer several advantages, including low equivalent weight, higher water solubility, steep positive temperature coefficients of solubility and safety in foods and pharmaceuticals (Deblay, 1995; Michaels, 1998). Typical transmembrane pressures encountered in osmotic distillation are in the range of 140 kPa (Hogan et al., 1998). Mass transfer aspects The water transport in the OD process is related to the driving force, represented by the vapour pressure difference at both liquid–vapour interfaces of the membrane (DPwm ), by the following equation:
Jw = Km (Pwf – Pws)
[9.1]
where Jw is the transmembrane flux, Km the membrane mass transfer coefficient, Pwf and Pws are the water vapour pressures of the feed and stripping solution at the membrane surface.
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Advances in membrane-based concentration 253 The water vapour pressures of the feed and stripping solutions at the fluid membrane interfaces are calculated as:
* Pwf = Pwf awf
[9.2]
* Pws = Pws aws
[9.3]
* * and Pws where Pwf are the pure water vapour pressures of the feed and stripping solution at the interface, respectively; awf and aws are the water activities of the feed and stripping solution at the interface, respectively. If the temperature is the same on both sides of the membrane equation [9.1] can be written as:
J w = K m Pw* (awf – aws )
[9.4]
The driving force depends on the solute concentration as well as on the temperature conditions prevailing at the vapour–liquid interfaces. A more detailed representation referring to the bulk conditions of both compartments is given by integrating the various mass transfer resistances:
lw = KD Pwb
[9.5]
where
1 1ˆ Ê1 K =Á + + ˜ K K K Ë f m s¯
–1
is defined as the overall mass transfer coefficient which accounts for the resistances opposed by the feed solution (1/Kf), the membrane (1/Km) and the stripping solution (1/Ks). If the concentration polarization is negligible, mass transfer depends only on membrane resistance and consequently:
K = K m
[9.6]
As in other membrane processes, OD is affected by concentration polarization. Owing to the water transport across the membrane, concentration in the bulk phase and, consequently, the water activity, differs from the composition near the membrane interface. As a result, the driving force for mass transfer and water flux are reduced (Ravindra Babu et al., 2006). Independently of the type of driving force (temperature difference, activity difference or both), the water vapour transfer mechanism in the membrane pores can be estimated on the basis of the molecular diffusion model, the Knudsen diffusion model or by a combination of both mechanisms (Geankoplis, 1993). If the mean free path of gas molecules is significantly greater than the membrane pore size, diffusing molecules collide frequently with pore
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254 Separation, extraction and concentration processes walls rather than with gas molecules present in the pores. In this instance the molecular transport occurs by Knudsen diffusion: 2e r K mK = M RT 3 cd
8RT pM
[9.7]
where r is the pore radius, e the membrane porosity, c the tortuosity factor, d the membrane thickness, M the molecular weight, R the gas constant and T the temperature. For pore size greater than the mean molecular free path of water vapour, the molecular diffusion is the controlling mechanism and the water mass flux can be written as: De P K mM = M RT cd Plm
[9.8]
where D is the diffusion coefficient and P the pressure. The Knudsen number (Kn) can be used to determine which of the two diffusion models is predominant. It compares the mean molecular free path (l) with the mean pore diameter of the membrane and is defined as: Kn = l 2p
[9.9]
where
l=
kBT P 2p s 2
[9.10]
in which kB is the Boltzmann constant, s is the mean collision diameter of the molecule and r the pore radius. For small pore size, Kn ≥ 10, collisions with pore walls are frequent therefore Knudsen diffusion is the prevailing mechanism. With relatively large pores, Kn ≤ 0.01, more collisions between the gas molecules themselves occur and the molecular diffusion will be predominant. Both mechanisms coexist when Kn is between the two limit values (Courel et al., 2000a). The concentration polarization phenomenon can be reduced by optimization of process parameters as stirring or stripping solution properties. The liquid concentration difference between the membrane and the bulk phase determines a diffusive solute flow according to Fick’s law which counterbalances the convective flow in the opposite direction. A mass balance across the boundary layer of the stripping solution is given by:
ÈÊ C ˆ ˘ J v = kp ln ÍÁ b ˜ ˙ ÎË Cm ¯ ˚
with
kp =
Dsw d
[9.11]
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Advances in membrane-based concentration 255 where Jv is the volume flux, Dsw the diffusion coefficient of solute by the membrane, kp the liquid mass transfer coefficient, Cm the salt molar concentration in the membrane, Cb the salt molar concentration in the bulk and d the boundary layer thickness. The liquid mass transfer coefficients depend on the solution physical properties and on the hydrodynamic properties of the system. These coefficients can be estimated by empirical correlations of dimensionless numbers namely Sherwood (Sh), Reynolds (Re) and Schmidt (Sc) numbers, equations [9.12] and [9.13].
Sh = a Reb Scg
[9.12]
with
Sh =
kdh , Dw
Re =
r udh , m
Sc =
m r Dw
[9.13]
where a, b and g are constants, k the feed or osmotic agent side liquid mass transfer coefficient, dh the hydraulic diameter, Dw the diffusion coefficient in water, r the solution density, m the solution dynamic viscosity and u the cross flow velocity of the fluid. Heat transfer OD is essentially a mass transfer process. The water transport through the membrane implies evaporation at the feed side and condensation at the stripping solution side. This phase change at the membrane walls generates a temperature difference across the membrane which reduces the vapour pressure difference decreasing the water transport. Under steady state conditions, the heat transfer equations can be written as:
Q = hf (Tf – Tfm) = hs (Tsm – Ts)
[9.14]
Q = Nw DHv – hm(Tsm – Tfm)
[9.15]
where Q is the heat flux, hf, hs and hm the heat transfer coefficients of the boundary layers (feed and stripping solution, respectively) and of the membrane, Tf and Tm the bulk temperatures of the feed and stripping solution, respectively, Tfm and Tsm the temperatures of the feed and stripping solution at the membrane interfaces, respectively, DHv the water latent heat of vaporization and Nw the molar vapour flux. The relation between heat flux and overall heat transfer coefficient during the OD process can be obtained by substituting Tsm and Tfm from equation [9.14] in equation [9.15], obtaining:
N DH Q = U ÈÍ w v – (Ts – Tf )˘˙ Î hm ˚
where U is the overall heat transfer coefficient given by: © Woodhead Publishing Limited, 2010
[9.16]
256 Separation, extraction and concentration processes
1 1ˆ Ê1 U=Á + + Ë hf hm hs ˜¯
–1
[9.17]
The thermal effect is considered negligible by most researchers. For example, Ravindra Babu et al. (2006, 2008) found that the contribution of concentration polarization on driving force reduction, during the concentration of pineapple and sweet-lime juices by OD, was prominent when compared with that of temperature polarization. On the other hand, the thermal effect owing to evaporation and condensation at both liquid–membrane interfaces was considered substantial by other authors (Courel et al., 2000a; Gostoli, 1999). In particular, the results of pure water OD experiments performed by Courel et al. (2000a) with commercial asymmetric porous membranes indicated that a high vapour flux of 12 kg m–2 h–1 generates a transmembrane temperature gradient of about 2 °C inducing a 30% driving force reduction. Osmotic distillation (OD) membranes and modules Hydrophobic polymers with low surface free energy are commonly used to produce membranes with pore sizes and pore size distribution suitable for OD applications. They include both polyolefins, such as polyethylene (PE) and polypropylene (PP), and perfluorocarbons such as polytetrafluoroethylene (PTFE) and polyvinylidene fluoride (PVDF). Membranes comprising these polymers, having a pore diameter ranging from 0.1 to 1 mm, allow the gas phase to be maintained in the membrane pores, a fundamental condition required to perform the OD process (Gryta, 2005). However, hydrophilic ceramic membranes can be modified by grafting on the surface molecules containing hydrophobic fluorocarbon chains like fluoroalkylsilanes or by coating the surface of alumina membranes with a thin lipid film. These membranes have been applied successfully in OD (Gabino et al., 2007; Romero et al. 2006). By referring to the configuration, hollow-fiber membranes, characterized by thin walls, are preferred for OD applications because they offer high surface/ volume ratios and do not require supports or spacers. Various parameters have to be considered in the selection of the membrane such as: pore size, porosity, conductivity and thickness. An increase in the pore size enhances the evaporation flux (flux is proportional to the radius); furthermore, membranes with large pore size exhibit a higher retention towards volatile organic flavour/ fragrance components (Barbe et al., 1998) than membranes with smaller pore size. These results can be attributed to differences in feed–membrane and stripper–membrane boundary layer resistances to organic volatiles transport owing to various degrees of liquid intrusion into the pores. Highly porous membranes are preferred because the evaporation flux is proportional to the porosity. Furthermore, an increase in membrane conductivity allows the heat gradient across the membrane to be minimized. Because the evaporation flux is proportional to the reciprocal of the pore length, the membrane thickness should be as low as possible. The thickness © Woodhead Publishing Limited, 2010
Advances in membrane-based concentration 257 is usually limited by the mechanical strength of the membrane: therefore thin membranes are supported by net, e.g. Gelman PTFE membranes. The overall thickness for OD membranes can vary from 80 to 250 mm, depending on the absence or presence of support. The risk of wetting of the hydrophobic membrane, with a consequent reduction in the evaporation flux and separation performance, is the main drawback of the osmotic distillation process. The reduction in the membrane thickness leads to a decrease in thermal resistance that facilitates the heat transfer from the brine to the feed causing a reduction in the temperature gradient (DT) and an increase in the partial pressure gradient (DP) across the membrane. However, thin hydrophobic membranes are more susceptible of wetting with a consequent loss of both evaporation flux and separation performance. If wetting occurs, the liquid can penetrate into pores in the membrane and, consequently, a liquid flux is added to the vapour flux and nonvolatile solutes diffuse across the membrane from one compartment to the other. Some fruit juices, especially citrus juices, contain peel oils and other highly hydrophobic compounds that promote wetting of OD membranes. Furthermore, cleaning solutions often contain surface-active agents that might also promote membrane wetting. Hydrophilic polymer films, which have sufficiently high intrinsic water permeability and are virtually impermeable to macrosolutes and colloids, are suitable as laminated membranes to prevent liquid intrusion without impeding vapour transport (Hogan et al., 1998). They include: esters and ethers of cellulose, crosslinked gelatin, chitin, agar, alginic acid, crosslinked polyacrylamide, crosslinked polyvinyl alcohol (PVA), polyhydroxy 2-ethyl methacrylate (PHEMA). Commercially available cellophane membranes can be suitable for this purpose; the hydrogel-film side of the laminate should be in contact with the solution to be concentrated (Michaels, 1999). Mansouri and Fane (1999) developed modified hydrophobic membranes for OD tolerant to oily feeds (i.e. limonene) by coating the feed side of commercial membranes with a thin layer of PVA. The uncoated membranes were rapidly wetted out by the oily feed even for low concentrations of oil (limonene) dispersed in water; vice versa, the coated membranes were stable for concentrations up to 1 wt.% limonene solution for periods up to 24 h. Commercial membranes for OD applications (termed TF200, TF450 and TF1000) manufactured by Pall-Gelman (East Hills, NY, USA) are characterized by a thin PTFE layer supported by a polypropylene (PP) net. In these asymmetric membranes, the top layer offers a resistance to gas transfer, whereas the membrane support offers an additional resistance to water transfer in the liquid form. The mass transfer resistance in the vapour phase is about 40–70% of the total resistance. The resistance of diluted brine entrapped in the PP support can cover up to 30% of the total resistance and the diluted brine boundary layer up to 60% indicating the sensitivity of the OD system to concentration polarization phenomena (Courel et al., 2000a). The best-known module designed for OD is the Liqui-Cel® Extra-Flow membrane contactor (Membrana-Charlotte, North Carolina, USA) containing
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258 Separation, extraction and concentration processes microporous PP hollow fibre membranes approximately 300 mm in external diameter with a wall thickness of about 40 mm; they have a mean pore diameter of about 30 nm and a porosity of about 40%. The fibres are potted into a polyethylene tubesheet and the shell casing is PP, PVDF or 316L stainless steel (Gabelman and Hwang, 1999; Hogan et al., 1998; Sirkar, 1997). The smallest modules are 2.5 ins in diameter with a membrane surface area of 1.4 m2, whereas the largest are 10 ins in diameter and offer a contact area of 130 m2. Commercial membranes commonly used in OD are summarized in Table 9.2. Effect of operating conditions on the OD flux The performance of the OD process in terms of water vapour transport is influenced profoundly by the operating conditions. Feed concentration inversely affected the performance of the OD process. Courel et al. (2000b) observed a decrease in evaporation flux (from 10.3 to 1.1 kg m–2 h–1) when sugar solutions of increasing sucrose content, from 0 to 65% w/w, were dehydrated at 25 °C by using stripping solution of 45.5 w/w% initial CaCl2 content. A transmembrane flux decay by increasing the feed concentration was observed also by Ravindra Babu et al. (2006) in the concentration of sweetlime juice and phycocianin solution by OD. The viscosity of the feed solution increases exponentially with the solute content whereas the diffusion coefficient strongly decreases relative to pure water. The increasing viscosity results in an increase of the concentration polarization effect reducing the driving force and, consequently, the flux rate (Bui et al., 2005). Sheng et al. (1991) studied the effect of operating conditions on the OD flux during the concentration of apple, orange and grape juice through a PTFE membrane with a pore size of 0.2 mm and an overall thickness of 100 mm. The osmotic pressure difference D p between the aqueous streams strongly affected the transmembrane vapour flux; in particular, a 33% decrease in D p determined a five-fold decline in OD flux. The OD flux is significantly affected also by the solute content of the stripping solution. Courel et al. (2000b) observed a 64% vapour flux decay for a 30% mass fraction reduction of the brine solution in extracting pure water at 25 °C by using CaCl2 as extractant. Ravindra Babu et al. (2006) observed a similar trend in the concentration of phycocianin solution and sweet-lime juice. These results may be explained by assuming the strong dependence of the water activity of the stripping solution on salt content. Although a vapour flux improvement is expected when the salt content is reduced (since the density and viscosity of the brine tend to decrease reducing the mass transfer resistance in the salt solution) this improvement is in fact masked by the activity effect. The OD flux is also differently affected by the type of osmotic agent. Calcium chloride produces higher transmembrane fluxes than sodium chloride © Woodhead Publishing Limited, 2010
Table 9.2 Typical membranes used in the OD process (adapted from Gryta, 2005) Manufacturer
Material
Configuration
Thickness (mm)
Porosity (%)
Average pore size (mm)
Durapore GVHP Durapore GVSP Durapore HVHP FHLP Accurel PP Q3/2 Accurel PP S6/2 Celgard 2500 Celgard 2400 Metricel SM35 TF200 TF450 TF1000 Gore-Tex 10387
Millipore Co. Millipore Co. Millipore Co. Millipore Co. Enka A.G. Enka A.G. Celgard LLC Celgard LLC Gelman USA Scimed Life Systems Inc. Pall-Gelman Pall-Gelman Pall-Gelman Gore & Associates
PVDF PVDF PVDF PTFE PP PP PP PP PP PDMS PTFE layer PTFE layer PTFE layer PTFE layer
Flat Flat Flat Flat Capillary Capillary Flat Flat Flat
125 108 125 175 200 400 25 25 90 300 178 178 178 8.5
70 80 75 70 70 70 50 37 55
0.20 0.22 0.45 0.25 0.2 0.2 0.07 0.05 0.1
80 80 80 78
0.2 0.45 1.0 0.2
supported supported supported supported
by by by by
PP PP PP PP
net net net net
Flat Flat Flat Flat
Advances in membrane-based concentration 259
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Type
260 Separation, extraction and concentration processes in the concentration of pineapple juice by OD (Nagaraj et al., 2006). This is mainly because of the higher osmotic activity of calcium chloride that results in a higher vapour pressure gradient across the membrane. Aqueous solutions of propylene glycol, glycerol and glycerol–salt mixtures were investigated as an alternative to calcium chloride in order to overcome the problem of corrosion and scaling associated with the use of brines (Celere and Gostoli, 2004). Propylene glycol and glycerol solutions (70–75 wt.%) were less effective than highly concentrated CaCl2 and exhibited a similar extractive power. However, propylene glycol cannot be recommended as an extractant in juice concentration owing to its low penetration pressure through the membrane pores and the not negligible volatility. Ternary mixtures water–glycerol–NaCl are characterized by lower viscosities than glycerol alone, and offer similar fluxes. The transmembrane flux in OD increases with an increase in the flow rate of the osmotic agent (Courel et al., 2000b; Nagaraj et al., 2006). This can be attributed to a stronger shear stress along the condensation side of the membrane leading to a reduction in the hydrodynamic boundary layer thickness and, consequently, to the concentration polarization effect. Ravindra Babu et al. (2008) observed an 8% increase in transmembrane flux when the flow rate of pineapple juice at about 12 °Brix was increased from 25 mL min–1 to 100 mL min–1; the increase in transmembrane flux can be explained by assuming a reduction in the concentration polarization effect on the feed side. The increase in flux, however, was more prominent (about 20%) by increasing the osmotic agent velocity. This phenomenon can be attributed to a lower concentration polarization on the feed side than on the brine side. A minor role of feed velocity on reducing the polarization problem in OD was observed also by Bui and Nguyen (2006) in the concentration of aqueous glucose solutions. They assumed that the heat transfer in OD is only a little involved or not involved at all; on the contrary, feed flow rate is essential to maintain the temperature gradient across the membrane in MD. The evaporation flux in OD is also affected by the feed temperature. Courel et al. (2000b) studied the influence of the temperature under isothermal conditions on the evaporation fluxes of pure water and sugar solutions of 35–65 w/w% sucrose content at 20–35 °C. A calcium chloride solution with an initial concentration of 45 w/w% was used as stripping solution. Evaporation fluxes ranged from 0.5 kg m–2 h–1 for a sugar solution of 65 w/w% (at 20 °C) to 12.8 kg m–2 h–1 for pure water (at 35 °C). The extent of mass transfer increase depended on the solute content: in the range 20–35 °C the evaporation flux increased by 120%, for a 35 w/w% sucrose solution, and only by 32% at 60 w/w% solute content. This phenomenon can be explained by assuming an exponential type relation between the vapour pressure difference across the membrane and the temperature according to Clapeyron’s law. Moreover, an increase in temperature results in a decrease in the feed and brine viscosities and an increase in the solute diffusion
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Advances in membrane-based concentration 261 coefficient: the higher the solute content, the stronger the thermal effect. Bui and Nguyen (2006) reported a 200% increase in the evaporation flux for a feed temperature increase of 20 degrees in the concentration of 40 w/w% and 50 w/w% aqueous glucose solutions by means of PVDF hollow fibres. OD applications Because OD can be operated at room temperature and atmospheric pressure, with minimal thermal and mechanical damage of the solutes, it has been successfully applied to the concentration of liquid foods such as milk, fruit and vegetable juices, instant coffee, tea and pharmaceuticals. The low temperature employed avoids the chemical or enzymatic reactions associated with heat treatment. The low operating pressure results in lower energy consumption and capital investment, reduced fouling phenomena and the possibility of using membranes characterized by lower mechanical resistance than those in pressure-driven membrane processes (Courel et al., 2001). OD is particularly suited for fruit juice concentration, producing concentrated juices with a quality and composition very close to fresh ones and higher than conventional products obtained by thermal evaporation (Cassano and Drioli, 2007; Jiao et al., 2004). Plate-and-frame modules with a net-shaped spacer on the extract side and a smooth juice side path were developed for the concentration of whole juice with a high pulp content (Cervellati et al., 1998). This configuration, even if not completely efficient in terms of mass transfer, allows unclarified juices to be processed. Helically wound hollow-fibre modules offer an improvement in the hydrodynamic conditions on the shell-side compared with axial flow modules; consequently, higher concentration of solutes and higher evaporation fluxes can be obtained when viscous feeds are processed (Costello et al., 1997). Rodrigues et al. (2004) evaluated the performance of the OD and RO processes in the concentration of camu-camu juice. RO allowed higher fluxes to be reached (50 kg m–2 h–1) than OD, but lower concentration of soluble solids (25 °Brix). OD allowed the juice to be concentrated up to 63 °Brix with evaporation flux values of 10 kg m–2 h–1. A pre-treatment of the whole juice by UF or MF allows evaporation fluxes to be improved in OD. However, the performance of the OD process is affected by pore diameters of UF membranes. Bailey et al. (2000) observed higher evaporation fluxes when Gordo grape juice was preliminarily ultrafiltered with membranes having pore diameters of 0.1 mm or less. UF membranes with a nominal pore diameter of 0.5 mm did not improve the OD fluxes. UF pretreatment results also in a small increase in juice surface tension with a consequent reduction in the tendency for membrane wet-out to occur. Several kinds of clarified fruit juices (orange, apple, grape, cactus pear, kiwi, passion fruit, and pineapple) were concentrated by OD (Alves and Coelhoso, 2006; Cassano and Drioli, 2007; Cassano et al., 2003, 2004, 2006,
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262 Separation, extraction and concentration processes 2007; Durham and Nguyen, 1994; Galaverna et al., 2008; Hongvaleerat et al., 2008; Jiao et al., 1991; Koroknai et al., 2006a, 2006b; Rektor et al., 2006; Rodrigues et al., 2004; Vaillant et al., 2001, 2005). An integrated membrane process was developed by Cisse et al. (2005) to produce concentrated orange juice. The process included a MF pre-treatment followed by OD. The quality of the concentrated juice was similar to the freshly squeezed one. Cassano et al. (2003, 2004, 2006 and 2007) proposed integrated membrane processes including UF, RO and OD to produce concentrated fruit juices such as kiwifruit, orange, lemon and cactus pear juice. Depectinized juices were clarified by UF and then optionally preconcentrated by RO. The OD process was employed as a concentration step to produce concentrated juices with final total soluble solids (TSS) up to 63–65 °Brix. The proposed process employs partial batch recycle on the feed side of the OD step in order to minimize feed viscosity changes through the membrane contactors and continuous counter current recycle plus evaporative reconcentration of the brine strip. The residual fibrous phase coming from the UF treatment (retentate) could be submitted to a stabilizing treatment (pasteurization, ohmic heating, high pressures, electric fields at high voltage) and successively added to the final OD concentrate for the preparation of fibre enriched beverages (Fig. 9.3). The total antioxidant activity (TAA) of the clarified or pre-concentrated juice was kept constant during the OD process independently of the achieved degree of TSS. Both aroma and colour were similar to those of the fresh juice. Ascorbic acid and health-promoting substances were also well preserved. On the contrary, a larger decrease in content of antioxidant compounds was detected in thermally evaporated juices with the same TSS content. In the concentrated orange juice, anthocyanins and hydroxycinnamates (particularly ferulic and p-coumaric acid) underwent a reduction of 36 and 55%, respectively; for the ascorbic acid and flavonones removals were in the order of 30% and 23%, respectively (Galaverna et al., 2008). Cassano and Drioli (2007) also evaluated the quality of kiwifruit juice concentrated by OD and thermal evaporation. The juice concentrated by thermal evaporation at 65 °Brix showed an 87% reduction of the ascorbic acid compared with the clarified juice. The TAA was reduced by about 50% independently of the TSS content achieved. On the contrary, the juice concentrated by OD retained almost all the TAA and the ascorbic acid content of the clarified juice. Blackcurrant juices concentrated by OD, with a final TSS content of 63–72 °Brix, showed a colour intensity, transparent ability and acidic flavour intensity similar to those of the raw juice (Kozák et al., 2008). A hybrid plant consisting of UF and RO pretreatment stages and an OD section (containing two 19.2 m2 Liqui-Cel membrane modules) for the concentration of fruit and vegetables juices was developed by Zenon Environmental (Burlington, Ont.). The plant was designed to produce concentrated juice at 65–70 °Brix at an average flow rate of 50 L h–1. It
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Fruit juice 10–11 °Brix
Reconcentrated brine
Feed reservoir Evaporator
Condenser
UF
Condensate
Stripping solution reservoir
OD Pulp
Preconcentrated juice 25–26 °Brix
RO
Water
Pasteurization
Pasteurized pulp
Reconstituted juice Concentrated juice 64–65 °Brix Water
Fig. 9.3 Integrated membrane process for the production of concentrated fruit juices (dashed lines refer to the production of the reconstituted juice).
Advances in membrane-based concentration 263
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Clarified juice
264 Separation, extraction and concentration processes was developed primarily for the concentration of grape juices destined for the production of high-quality vintage varietal wines such as Chardonnay, Cabernet Sauvignon and Merlot. Concentrates can be stored for long periods without deterioration and can be used as blending stocks to adjust the sugar content of freshly harvested grapes to minimize variations in alcohol content of the resulting wine between vintages. Additionally, the concentrated juice can be shipped over long distances and used to produce high-priced varietal wines starting from local grapes (Hogan et al., 1998). Evaporation fluxes between 7 and 10 kg m–2 h–1 were obtained by Hongvaleerat et al. (2008) in the concentration of clarified pineapple juice. These values were higher than those obtained with the single strength juice owing to the complete removal of suspended solids in the clarification step. At low TSS the evaporation flux decay during the OD process is more attributable to the dilution of the stripping solution whereas, at higher feed concentrations, the evaporation flux depends mainly on juice viscosity and, consequently, on juice concentration (Cassano and Drioli, 2007; Courel et al., 2000a; Vaillant et al., 2001). Evaporation fluxes in OD can be improved by application of acoustic fields. Narayan et al. (2002) observed an enhancement of the evaporation flux from 0.81 to 0.94 L m–2 h–1 during the OD of sugarcane juice with CaCl2 as stripping solution when ultrasounds were applied to the membrane cell. Concentration and temperature polarizations in OD contribute up to 18% to the flux reduction (Bui et al., 2005). However, concentration polarization gave a larger contribution to the flux reduction than temperature polarization. Furthermore, the flux reduction owing to polarization was smaller on the feed side than on the brine side. Hydrophobic membranes with relatively large pore sizes showed higher organic volatiles retention per unit water removal than those with smaller pores when model aroma aqueous solutions were processed by OD (Barbe et al., 1998). Pores with larger diameter at the membrane surface allow greater intrusion of the feed and brine solutions with a consequent increase in the thickness and resistance of the boundary layer at the pore entrance. Consequently, membranes with large surface pore diameters are preferred for OD applications in which retention of volatile flavour/fragrance components are required. Shaw et al. (2001) investigated the retention of flavours in concentrated orange and passion fruit juices prepared by OD. Losses of volatile components between 32 and 39% were observed in the concentrated products. The transfer of volatiles depends on the nature of compounds and on the operating conditions. A significant reduction in the transfer of juice volatiles was obtained by decreasing the circulation velocity or the temperature (Ali et al., 2002). The management of the diluted brine step is one of the drawbacks associated with the commercial application of the OD in fruit juice processing. Although the regeneration of exhausted brines can be realized by thermal evaporation, this operation is expensive owing to corrosion and scaling phenomena. Solar
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Advances in membrane-based concentration 265 ponding, reverse osmosis and pervaporation were proposed as alternatives to re-concentrate the diluted brine solutions (Thompson, 1991). Electrodialysis was also suggested for the regeneration of NaCl brines (Petrotos and Lazarides, 2001). The OD process is also a valid approach for producing low-alcohol beverages because it permits a selective removal of single volatile solutes from aqueous solutions (e.g. ethanol from wine and other ferments) by using water as a stripping solvent. The removal of organic solvents after the extraction of intracellular products (antibiotics, hormones and biologicals) from fermentation broths is a necessity in the drug industry. This operation has to be carried out at low temperature in order to preserve the product from thermal degradation and can be performed by OD in which water is employed as stripping solution (Hogan et al., 1998). The concentration of pharmaceuticals and biological compounds such as vaccines, hormones, recombinant proteins, enzymes, antibiotics, fungicides and nucleic acids is another potential use of OD. These products have to be isolated or recovered in the dry state in order to maintain their activity and shelf stability for a long time; OD can be used as a concentration step, eventually after a pre-concentration unit (by RO or NF), in order to obtain concentrates from which bioactive compounds can be more easily recovered (e.g. by crystallization or extraction). OD can be used also as a pre-concentration step for the lyophilization, reducing the water removal load during freeze drying. 9.4.2 Membrane distillation Process fundamentals Similarly to the OD process, in MD the water vapour transfer is promoted by a vapour pressure difference between two sides of a microporous hydrophobic membrane; however, in MD the physical origin of the vapour pressure difference is a temperature gradient rather than a concentration gradient: the feed is maintained at high temperature while cold water is used as a stripping permeate. Therefore, membrane distillation is a thermaldriven process. The mass transfer in MD can be described as a three-phase sequence: (1) formation of a vapour gap at the warm solution–membrane interface; (2) transport of the vapour phase through membrane pores; (3) its condensation at the cold side membrane–solution interface. Theoretical aspects and potential applications of the MD process have been extensively studied (Bandini et al., 1991; Calabrò et al., 1994; El-Bourawi et al., 2006; Lawson and Lloyd, 1997; Schofield et al., 1987; Tomaszewska, 2000a). The process takes place at atmospheric pressure and at temperatures that may be much lower than the boiling point of the treated solutions. Consequently, it can be used to concentrate solutes sensitive to high
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266 Separation, extraction and concentration processes temperatures such as fruit juices and pharmaceuticals. Feed temperatures in MD are typically 60–90 °C, although temperatures as low as 30 °C have been used. Operating pressures are generally of about 0–100 kPa, hence much lower than conventional pressure-driven membrane processes such as RO. Consequently, lower equipment costs and increased process safety can be achieved. Furthermore, the mechanical resistance of the membrane is greatly reduced. Because MD operates on the principles of vapour–liquid equilibrium, another advantage over traditional pressure-driven membrane processes is represented by its high rejection (theoretically 100%) towards ions, macromolecules, colloids, cells and other nonvolatile compounds (Lawson and Lloyd, 1997). Finally, vapour spaces can be reduced compared with conventional distillation processes. As in the OD process, the main drawback of MD is the risk of wetting of the hydrophobic membrane with a consequent reduction in the evaporation flux and separation performance. Therefore, the process solutions must be aqueous and sufficiently dilute. This limits MD to applications such as desalination, removal of trace volatile compounds from wastewater and concentration of nonvolatile aqueous solutions. Membrane distillation (MD) configurations As depicted in Figure 9.4, the MD process can be realized according to four types of configuration. In the direct contact membrane distillation (DCMD), the membrane separates the hot feed from the cold distillate. In this instance, the vapour pressure gradient, which results from the transmembrane temperature difference, is the driving force of the mass transport across the membrane. Volatile molecules evaporate at the hot liquid/vapour interface, cross the membrane pores in the vapour phase, and condense on the cold liquid/vapour interface inside the membrane module (Gryta, 2002; Lawson and Lloyd, 1997; Mengual and Peña, 1997; Tomaszewska, 2000b). DCMD is usually employed for applications in which water is the major fluxing component, such as desalination or concentration of aqueous solutions. In air gap membrane distillation (AGMD), a condensing surface is separated from the membrane by an air gap. Volatile molecules cross both the membrane pores and the air gap and finally condense over a cold surface inside the membrane module. In the sweeping gas membrane distillation (SGMD), a cold inert, gas sweeps the permeate side of the membrane carrying the volatile molecules. In this instance, condensation occurs outside the membrane module (Khayet et al., 2000). In the vacuum membrane distillation (VMD), vacuum is applied on the permeate side of the membrane by means of a vacuum pump and, similarly to SGMD, condensation takes place outside the membrane module (Bandini et al., 1992).
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Feed
Retentate
Feed
Retentate Membrane
Membrane
Permeate
Permeate in
Cooling water out
Air gap
Condensing plate
(a)
Feed
(b)
Retentate
Feed
Retentate Membrane
Membrane
Sweep gas in Condenser
Permeate
Sweep gas out
(c)
Cooling water in
Condenser
Permeate
Vacuum
(d)
Fig. 9.4 Membrane distillation configurations: (a) DCMD; (b) AGMD; (c) SGMD; (d) VMD.
Advances in membrane-based concentration 267
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Permeate out
268 Separation, extraction and concentration processes Mass transfer and polarization phenomena In a similar way to that of the OD process, classical models of gas diffusion in porous membranes, such as molecular diffusion or Knudsen diffusion, can be used to describe the mass transfer across MD membranes (Courel et al., 2000a; Gryta, 2005). These models suggest a linear relationship between the volume flux per unit surface area of the membrane and the transmembrane water vapour pressure difference at each interface, which depends on composition and temperature of both compartments. The osmotic and thermal contributions can operate either in a synergistic or in an antagonistic way (in which one of them prevails on the other) (Godino et al., 1995). The relationship between the permeate flux (Jw) and the driving force (DP) which causes the mass transfer across the membrane pores is given by:
Jw = Kw DP
[9.18]
where Km is the membrane mass transfer. The driving force in MD is a partial pressure gradient in the vapour phase given by:
DP = Pwf – Pwd
[9.19]
where Pwf and Pwd are water vapour pressures at the feed/membrane and feed/distillate interface, respectively. The mass transfer through the membrane, as in other membrane processes, is caused by the chemical potential difference (Dm) on both sides of the membrane (distillate and feed sides), which, for a ith component is given by:
D mi = D mid – D mif = RT ln
apid a = RT ln id apif ai
[9.20]
where apid and aid are the activity of ith component in the vapour and liquid if of ith component in phase on the distillate side, apif and aif the activity the vapour and liquid phase on the feed side, R the gas constant and T the temperature. The driving force expressed in equation [9.19] requires a knowledge of both temperature and solute composition at the vapour–liquid interfaces. Because the interfacial conditions are not always accessible, the water transport, referred to the bulk conditions, is often described as:
1 1ˆ Ê1 b b J w = K (Pwf – Pwd )=Á + + Ë K f K m K d ˜¯
–1
DP Pwb
[9.21]
where the overall mass transfer coefficient (K) is given by the series of water transport resistances from the bulk of the dilute solution towards the evaporation surface (1/Kf), water vapour transport through the membrane
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Advances in membrane-based concentration 269 (1/Km) and from the condensation surface to the bulk of the distillate side (1/Kd). The overall mass transfer coefficient depends on the membrane morphology, module design, solution concentration, temperature and hydrodynamic conditions (Gryta et al., 2005). When the temperatures in the layers near the membrane differ from those measured in the bulk of the feed and permeate side, a temperature polarization phenomenon occurs. However, temperature profiles formed in the DCMD are different from those observed in OD. The temperature gradient results from the evaporation at the feed side and condensation at the permeate side even if the bulk temperatures of two liquids are equal, as in OD. However, the heat conduction from the brine side to the feed side decreases the temperature polarization phenomenon in OD. On the contrary, in the case of DCMD, heat and mass transfer flow through the membrane in the same direction (from the feed to the distillate side) increasing the temperature polarization effect (Gryta, 2005). In all MD variants, as in OD, only water passes through the membrane causing an increase in the solute concentration on the feed side and a decrease in the solute concentration on the distillate side relative to the bulk condition. This phenomenon, called concentration polarization, results in a driving force reduction across the membrane. Several authors assume that the concentration polarization is substantial in OD, whereas the performance of DCMD is affected mainly by the temperature polarization effect, except for the concentration of solutions with a solute content close to the saturated state (Gryta, 2002). Alves and Coelhoso (2006) compared the performance of OD and MD in the concentration of sucrose solution (used as a model fruit juice), in terms of water flux and aroma retention. At the same applied overall driving force, water fluxes in MD were less than half of those observed in OD owing to temperature polarization effects. Moreover a higher retention of aroma compounds was observed for the OD process. MD membranes MD membranes must be porous and hydrophobic with good thermal stability and excellent chemical resistance to feed solutions. They must be characterized by high liquid entry pressure which is defined as the minimum hydrostatic pressure that must be applied to the feed liquid solution before it overcomes the hydrophobic forces and penetrates into the membrane pores. Membranes with high liquid entry pressure can be obtained by using materials with high hydrophobicity and small maximum pore size. The most suitable materials for hydrophobic MD membranes include PVDF, PP, PE and PTFE. These membranes are available in capillary or flatsheet configurations (Ding et al., 2002). PVDF appears of particular interest in MD processes because of its high melting point and good temperature resistance; its resistance to oxidation and gamma radiation, to solvents and to abrasion are also of interest. Hydrophilic membranes, such as cellulose
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270 Separation, extraction and concentration processes acetate membranes, can be treated in order to make their surfaces hydrophobic and, consequently, suitable for MD applications (Wu et al., 1992). Composite MD membranes consisting of hydrophobic and hydrophilic layers or a hydrophobic layer sandwiched between two hydrophilic layers have been also realized (Cheng et al., 1982). The size of micropores can range between 100 nm and 1.0 mm. In particular, the pores should be large enough to facilitate the required flux but also small enough in order to prevent liquid penetration through the membrane. For cylindrical pores, the maximum critical pore size at which the liquid penetrates the microporous phase is defined by the Laplace equation:
P = 2g
cos q r
[9.22]
where g is the surface tension of the liquid, q the contact angle between the liquid and the membrane, r the radius of the pore and P the applied pressure. For a given pore size a critical pressure Pc exists. For applied hydrostatic pressures higher than the Pc values, the liquid phase is transported across the membrane. Microporosity can be induced by mechanical stretching or by thermal phase separation technique. The relationship between the transmembrane flux and membrane characteristic parameters is given as:
N µ
ra e dm c
[9.23]
where N is the molar flux, r the mean pore size of the membrane pores, a a factor whose value equals 1 or 2 for Knudsen diffusion and viscous fluxes, respectively, dm the membrane thickness, e the membrane porosity and c the membrane tortuosity. According to equation [9.23], the MD flux increases with the increase in pore size. However, in order to avoid membrane pores wettability, the pore size should be as small as possible. Consequently, an optimum value for the pore size has to be determined for each MD application on the basis of the type of the feed solution to be treated. Schneider et al. (1988) recommended a maximum pore radius of 0.5–0.6 mm in order to avoid membrane wetting owing to fluctuations in process pressure and temperature. As for other membrane processes, the permeate flux in MD is inversely proportional to the membrane thickness: consequently thinner membranes produce higher fluxes. On the contrary, the membrane should be as thick as possible in order to prevent heat loss by conduction through the membrane matrix (Schofield et al., 1987). An optimal membrane thickness, considering the thermal conductivity of commercial membranes, should be within the range of 30–60 mm (Laganà et al., 2000). The porosity is the most influential factor affecting mass transfer rate in MD: membranes with higher porosity produce higher evaporation fluxes (Schneider et al., 1988). Membranes
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Advances in membrane-based concentration 271 with high porosity also allows the amount of heat lost by conduction to be reduced, because the heat transfer coefficient of the gas entrapped within the membrane pores is generally an order of magnitude smaller than that of the membrane material. MD membrane modules can be realized with flat-sheet membranes, as plate and frame and spiral-wound configurations, or with capillary membranes in tubular configuration. The design of the MD modules provides a high feed flow rate, to increase turbulence of the feed solution, low pressure drop and high packing density; moreover, because the MD is a nonisothermal process, good heat recovery and thermal stability is also guaranteed. Laboratory-scale modules are usually realized with flat-sheet membranes that are much more versatile than capillary membranes. They can be easily removed from their modules for cleaning, examination or replacement. Consequently, the same membrane module can be used to test different MD membranes. On the other hand, tubular membrane modules are preferred for commercial applications because they do not require a support and allow high membrane surface area/ module volume ratios to be realized. Effect of operating parameters on MD fluxes Evaporation fluxes in MD increase with an increase in the feed temperature. This is because the exponential increase of the vapour pressure of the feed solution with temperature increases the vapour pressure difference across the membrane and, hence, the driving force of the process. In DCMD applications the effect of decreasing the permeate temperature is to increase the evaporation flux (Lawson and Lloyd, 1999). A reduction in the flow rate implies a reduction in the Reynolds number and, consequently, in the transport coefficients. When the mean temperature is kept constant, the permeate flux increases linearly with the temperature difference. On the other hand, when the temperature difference is fixed, the permeate flux increases exponentially with the mean temperature (Mengual and Peña, 1997). When non-volatile solutes are considered, permeate fluxes in all MD configurations decrease by increasing the feed inlet concentration: this phenomenon can be attributed to the reduction of the driving force owing to the decrease in the vapour pressure of the feed solution and to the exponential increase in the viscosity of the feed solution. At high concentration ratios, MD fluxes are higher than those observed in other pressure-driven membrane processes (Cath et al., 2004). Permeate fluxes generally increase by increasing the solute concentration when aqueous solutions containing volatile components (such as alcohols) are processed. This can be explained by assuming that an increase in the volatile compound concentration in the feed side is associated with an increase in its transmembrane partial pressure. The increase in the permeate flow velocity causes an increase in the permeate flux in MD process. The increase in permeate flow velocity increases the heat transfer in the permeate side of the membrane reducing the temperature and concentration polarization phenomena. Consequently, the temperature
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272 Separation, extraction and concentration processes at the membrane surface approaches the temperature in the bulk permeate side increasing the driving force and the permeate flux. Finally, a linear increase in the MD flux with the transmembrane vapour pressure difference is typically observed in all MD configurations considered above. MD applications MD configurations previously described can be applied in different areas of interest for the separation of both nonvolatile (ions, colloids, macromolecules) and volatile (benzene, chloroform, trichloroethylene) compounds from water, the extraction of organic compounds such as alcohols from diluted aqueous solutions, the production of distilled water and the concentration of aqueous solutions. Desalination and pure water production from brackish water is the bestknown MD application (Banat and Simandl, 1998; Khayet et al., 2003, 2005). The concentration of radioactive solutions and wastewater treatment in the nuclear industry is another area under investigation (Khayet et al., 2006). In the chemical industry, the separation of azeotropic aqueous mixtures such as alcohol–water mixtures, the removal of volatile organic compounds from water and the concentration of acids, such as sulfuric, hydrochloric and nitric acid, can be achieved by MD (Banat and Simandl, 2000; Duan et al., 2001; Garcia-Payo et al., 2002; Tomaszewska, 2000b; Tomaszewska et al., 1995). Removal of dyes and wastewater treatment in the textile industry have also been reported (Banat et al., 2005; Calabrò et al., 1991). In the pharmaceutical and biomedical areas MD proved attractive for the removal of water from blood and protein solutions and in the wastewater treatment (Ortiz de Zárate et al., 1998; Sakai et al., 1986). MD can be successfully applied in the food industry and in areas where high temperature applications lead to the degradation of process fluids such as concentration of fruit juices and milk processing. A blackcurrant juice concentrate with a high TSS content was produced by MD without degradation of its valuable substances maintaining the juice side temperature at 26 °C and the water side temperature at 11 °C (DT = 15 °C) (Kozák et al., 2009). An increase of a few degrees Celcius in the driving force (DT = 19 °C) influenced significantly the evaporation flux and the operation time of the process. Gunko et al. (2006) used DCMD to investigate the concentration of apple juice. A TSS content of 50 °Brix was obtained when the permeate flux reached about 9 L m–2 h–1. Further concentration to 60–65 °Brix resulted in reduced productivity (up to 3 L m–2 h–1). Highly concentrated apple juices up to 64 °Brix were also produced by using polypropylene hollow fibre DCMD modules with trans-membrane fluxes of 1 kg m–2 h–1 (Curcio et al., 2000; Laganà et al., 2000). Flux rates were dependent essentially upon temperature polarization phenomena located mainly on the feed side, rather than concentration polarization. PVDF membranes used for the concentration of orange juice by MD
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Advances in membrane-based concentration 273 showed a very good retention of soluble solids, sugars and organic acids with rejection of sugars and organic acids equal to 100%. The colour and flavour of concentrated juice were satisfactory (Calabrò et al., 1994; Drioli et al., 1992). A protein concentrate containing at least 25% of protein was obtained by Christensen et al. (2006) in the processing of whey protein by DCMD as an alternative to thermal evaporation. The optimal temperature for the whey protein concentrate was 55 °C leading to less protein denaturation than evaporation and therefore a higher quality product. Bandini and Sarti (2002) studied the vacuum membrane distillation (VMD) for the concentration of grape must up to 50 °Brix. The process allowed production of juice concentrates that still retained interesting amounts of the aroma compounds. Bagger-Jørgensen et al. (2004) and Diban et al. (2009) evaluated the potential of VMD to recover aroma compounds from blackcurrant and pear juice, respectively. The highest values of enrichment factor (up to 15) for pear aroma compounds were obtained working at lower temperatures and higher downstream pressures. The highest concentration factors for the blackcurrant aroma compounds (from 21 to 31) were obtained at high feed flow rate (400 L h–1) and low temperatures (10 °C). At 5 vol.% feed volume reduction the recovery of highly volatile compounds was between 68 and 83 vol.% and between 32 and 38 vol.% for the hardly volatile compounds. Also in MD, an improvement in the evaporation flux in fruit juice concentration can be obtained when the juice is preliminarily submitted to a UF treatment. Drioli et al. (1992) found that the UF of the single-strength orange juice resulted in an increase in MD flux that remained almost constant during an approximately two-fold concentration. On the contrary, the MD flux of the unfiltered juice decreases steadily over the same concentration range. The improvement of the MD flux can be attributed to a reduction of juice viscosity as a result of pulp and pectin removal. PVDF membranes showed a very good retention of orange juice components such as total soluble solids, sugars and organic acids; on the other hand, the amount of ascorbic acid decreased by 42%, probably owing to its degradation associated with high temperature and oxidation. By referring to the retention of volatile organic flavour/fragrance components from liquid foods it has been found that membranes having an open fibrous structure rather than discrete pore offer the best volatiles retention for a given amount of removed water. 9.4.3 Coupled operation of osmotic distillation and membrane distillation Table 9.3 summarizes the main advantages and disadvantages of MD and OD processes. Because the water flux in OD is relatively low (usually between 0.07 and 7.2 kg m–2 h–1), the surface area available for flux should be increased. As described above, the thermal effect in OD reduces the driving
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274 Separation, extraction and concentration processes Table 9.3 Positive and negative aspects of MD and OD processes Advantages
Disadvantages
Low operating temperature
Low evaporative capacity with a long time of treatment Possibility of wetting of the hydrophobic membrane and consequent loss of flux and separation performance Production costs higher than those of thermal evaporation High cost of membrane replacement
Low operating pressure
No or less degradation of heat-sensitive compounds High TSS concentration in the concentrated products Modularity, easy scale-up Management of diluted brine solutions Possibility to treat solutions with high levels of suspended solids Possibility to concentrate several different products with the same unit No fouling problems Constant permeate flux in time Low investment cost
force of the water transport across the membrane. This phenomenon can be exploited to obtain a coupled process where the brine and feed solutions are thermostatically controlled at different temperatures: the osmotic solution on the cold side and the solution to be concentrated on the warm side. This coupled operation of MD and OD, referred to as membrane osmotic distillation (MOD), allows the water flux across the membrane to be enhanced (Wang et al., 2001). Although the aqueous solution is gently heated during the operation, this coupled method still works under mild conditions, because the temperature difference applied is lower than 15 °C (Bélafi-Bakó and Koroknai, 2006). The effect of varying experimental conditions (solute concentration, stirring rate, mean temperature and bulk temperature difference) on the water flux involved in a coupled process MD/OD was investigated by Godino et al. (1995). Koroknai et al. (2006a and 2006b) obtained concentrated fruit juices (apple, red- and blackcurrant, sour cherry and raspberry) at 60 °Brix in an operation time of 15–20 h, maintaining a temperature difference of 15 °C and using as an osmotic solution CaCl2 6M. In particular, the driving force of the process was greatly enhanced by decreasing the temperature of the osmotic solution, as in the range 15–25 °C only a slight drop in the water vapour pressure, for saturated CaCl2 solution, occurs. Consequently, maintaining the osmotic solution at room temperature, the energy consumption is minimized and the coupled process can be accelerated by increasing the temperature of the solution to be concentrated. Similarly, Bélafi-Bakó and Koroknai (2006) found that the coupled process is more effective than MD © Woodhead Publishing Limited, 2010
Process
Maximum achievable concentration (°Brix)
Product quality
Evaporation rate or flux
Possibility of treating Operating cost different products with the same installation
Thermal evaporation Cryoconcentration Reverse osmosis Direct osmosis Membrane distillation Osmotic distillation
65–75 30–50 25–30 50 60–70 60–70
Poor Very good Very good Good Good Very good
200–300 L h–1 – 5–10 L m–2 h–1 1–5 L m–2 h–1 1–10 L m–2 h–1 1–10 L m–2 h–1
No No No Yes Yes Yes
Moderate High High High High High
Capital investment
Energy consumption
Moderate Very high Moderate Moderate Low Low
Very high Very high High Low Low Low
Advances in membrane-based concentration 275
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Table 9.4 Key factors of conventional evaporation and membrane concentration techniques (adapted from Jariel et al., 1996)
276 Separation, extraction and concentration processes or OD alone. The use of short membrane modules in cascade series, with heat exchangers placed between them, was suggested in order to minimize heat losses.
9.5 Conclusions Key factors of conventional evaporation and membrane concentration techniques are summarized in Table 9.4. Several advantages of membranebased processes over conventional evaporation have been successfully demonstrated, including improved product quality, easy scale-up and low energy consumption. Low evaporation fluxes in OD and MD seem to be the main drawbacks of these processes when compared with RO and thermal evaporation. However, when the solution to be concentrated contain solutes sensitive to mechanical or thermal degradation there are serious limitations in using these technologies without significant deterioration in the quality. Currently membrane contactors are emerging technologies in the processing of foods, pharmaceuticals and beverages and will become breakthrough technologies when enhanced effectiveness is attained. Their integration with standard membrane operations is a valid approach for a sustainable industrial growth within the process intensification strategy. The aim of this strategy is to introduce in the productive cycles new technologies characterized by low encumbrance volume, advanced levels of automation capacity, modularity, remote control and reduced energy consumption. In order to improve the competitiveness of membrane contactors towards conventional technologies, efforts should be devoted to the development of new membranes characterized by high selectivity and stability for long-term applications, as well as to improvements in process engineering, including module and process design.
9.6 Nomenclature a a i api C d D h J k B k
water activity activity of the ith component in the liquid phase activity of the ith component in the vapour phase (fugacity) solute molar concentration (mol L–1) diameter (m) diffusion coefficient (m2 s–1) liquid heat transfer coefficient (W m–2 K–1) mass flux (kg m–2 h–1) or volume flux (m3 m–2 s–1) Boltzmann constant (1.3807 ¥ 10–23 J K–1) mass transfer coefficient (m s–1)
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Advances in membrane-based concentration 277 K Kn M N P P* Q r R Re Sc Sh T u U Greek g d e q DP Dm l m C r s
mass transfer coefficient (kg m–2 h–1 Pa–1) Knudsen number molecular weight (kg mol–1) molar flux (mol m–2 s–1) water vapour pressure (Pa) pure water vapour pressure (Pa) heat flux (W m–2) pore radius (m) gas constant (8314 J mol–1 K–1) Reynolds number Schmidt number Sherwood number temperature (°C, K) velocity of the fluid (m s–1) overall heat transfer coefficient (W m–2 K–1) symbols liquid surface tension (N m–1) thickness (m) porosity contact angle driving force chemical potential difference mean molecular free path (m) liquid dynamic viscosity (Pa s) tortuosity liquid density (kg m–3) mean collision diameter (m)
Subscripts b bulk d distillate f feed h hydraulic m membrane p pore or permeate s stripping solution or solute w water or vapour Superscripts b bulk location K Knudsen diffusion m membrane location M molecular diffusion
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278 Separation, extraction and concentration processes Abbreviations DO direct osmosis MD membrane distillation MF microfiltration OA osmotic agent OD osmotic distillation OMD osmotic membrane distillation PP polypropylene PTFE polytetrafluoroethylene PVDF polyvinylidene fluoride RO reverse osmosis TAA total antioxidant activity TSS total soluble solids UF ultrafiltration
9.7 References Aider M and de Halleux D (2008), ‘Production of concentrated cherry and apricot juice by cryoconcentration technology’, LWT-Food Sci Technol, 41, 1768–1775. Ali F, Dornier M, Duquenoy A and Reynes M (2002), ‘Transfer of volatiles through PTFE membrane during osmotic distillation’, Proceedings of the 2002 International Congress on Membrane and Membrane Processes, Toulouse, France. Alves VD and Coelhoso IM (2002), ‘Mass transfer in osmotic evaporation: effect of process parameters’, J Membrane Sci, 208, 171–179. Alves VD and Coelhoso IM (2006), ‘Orange juice concentration by osmotic evaporation and membrane distillation: a comparative study’, J Food Eng, 74, 125–133. Bagger-Jørgensen R, Meyer AS, Varming C and Jonsson G (2004), ‘Recovery of volatile aroma compounds from black currant juice by vacuum membrane distillation’, J Food Eng, 64, 23–31. Bailey AFG, Barbe AM, Hogan PA, Johnson RA and Sheng J (2000), ‘The effect of ultrafiltration on the subsequent concentration of grape juice by osmotic distillation’, J Membrane Sci, 164, 195–204. Banat FA and Simandl J (1998), ‘Desalination by membrane distillation: a parametric study’, Sep Sci Technol, 33, 201–226. Banat FA and Simandl J (2000), ‘Membrane distillation for propane removal from aqueous stream’, J Chem Technol Biotechnol, 75, 168–178. Banat FA, Al-Asheh S and Qtaishat M (2005), ‘Treatment of waters colored with methylene blue dye by vacuum membrane distillation’, Desalination, 174, 87–96. Bandini S, Gostoli C and Sarti GC (1991), ‘Role of mass and heat transfer in membrane distillation process’, Desalination, 81, 91–106. Bandini S, Gostoli C and Sarti GC (1992), ‘Separation efficiency in vacuum membrane distillation’, J Membrane Sci, 73, 39–52. Bandini S and Sarti GC (2002), ‘Concentration of must through vacuum membrane distillation’, Desalination, 149, 253–259. Barbe AM, Bartley JP, Jacobs AL and Johnson RA (1998), ‘Retention of volatile organic flavor/fragrance components in the concentration of liquid foods by osmotic distillation’, J Membrane Sci, 145, 67–75. Beaudry EG and Lampi KA (1990), ‘Membrane technology for direct osmosis concentration of fruit juices’, Food Technol, 44(6), 121.
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Advances in membrane-based concentration 279 Bélafi-Bakó K and Koroknai B (2006), ‘Enhanced water flux in fruit juice concentration: coupled operation of osmotic evaporation and membrane distillation’, J Membrane Sci, 269, 187–193. Bui AV, Nguyen HM and Muller J (2005), ‘Characterization of the polarizations in osmotic distillation of glucose solutions in hollow fibre module’, J Food Eng, 68, 391–402. Bui AV and Nguyen HM (2006), ‘The role of operating conditions in osmotic distillation and direct contact membrane distillation – a comparative study’, Int J Food Eng, 2(5) Art. 1, DOI: 10.2202/1556–3758.1171. Calabrò V, Drioli E and Matera F (1991), ‘Membrane distillation in the textile wastewater treatment’, Desalination, 83, 209–224. Calabrò V, Jiao B and Drioli E (1994), ‘Theoretical and experimental study on membrane distillation in the concentration of orange juice’, Ind Eng Chem Res, 33, 1803–1808. Cassano A, Conidi C, Timpone R, D’Avella M and Drioli E (2007), ‘A membrane-based process for the clarification and the concentration of the cactus pear juice’, J Food Eng, 80, 914–921. Cassano A, Drioli E, Galaverna G, Marchelli R, Di Silvestro G and Cagnasso P (2003), ‘Clarification and concentration of citrus and carrot juices by integrated membrane processes’, J Food Eng, 57, 153–163. Cassano A, Figoli A, Tagarelli A, Sindona G and Drioli E (2006), ‘Integrated membrane process for the production of highly nutritional kiwifruit juice’, Desalination, 189, 21–30. Cassano A, Jiao B and Drioli E (2004), ‘Production of concentrated kiwifruit juice by integrated membrane processes’, Food Res Int, 37, 139–148. Cassano A and Drioli E (2007), ‘Concentration of clarified kiwifruit juice by osmotic distillation’, J Food Eng, 79, 1397–1404. Cath TY, Adama VD and Childress AE (2004), ‘Experimental study of distillation using direct contact membrane distillation: a new approach to flux enhancement’, J Membrane Sci, 228, 5–16. Celere M and Gostoli C (2002), ‘The heat and mass transfer phenomena in osmotic membrane distillation’, Desalination, 147, 133–138. Celere M and Gostoli C (2004), ‘Osmotic distillation with propylene glycol, glycerol and glycerol–salt mixtures’, J Membrane Sci, 229, 159–170. Cervellati A, Zardi G and Gostoli C (1998), ‘Gas membrane extraction: a new technique for the production of high quality juices’, Fruit Process, 10, 417–421. Cheng DY and Wiersma SJ (1982), ‘Composite membrane for membrane distillation system’, US Patent 4,316,772. Cheryan M (1998), Ultrafiltration and microfiltration handbook, Technomic Publishing Co., Lancaster, PA. Christensen K, Andresen R, Tandskov I, Norddahl B and du Preez JH (2006), ‘Using direct contact membrane distillation for whey protein concentration’, Desalination, 200, 323–325. Cisse M, Vaillant F, Perez A, Dornier M and Reynes M (2005), ‘The quality of orange juice processed by coupling crossflow microfiltration and osmotic evaporation’, Int J Food Sci Technol, 40, 105–116. Collins AR and Harrington V (2002), ‘Antioxidants; not the only reason to eat fruit and vegetables’, Phytochem Revs, 1, 167–174. Costello AJ, Hogan PA and Fane AG (1997), ‘Performance of helically-wound hollow fibre modules and their application to isothermal membrane distillation’, Proceedings of Euromembrane ’97, 23–27 June, University of Twente, The Netherlands, 403–405. Courel M, Dornier M, Herry JM, Rios GM and Reynes M (2000b), ‘Effect of operating conditions on water transport during the concentration of sucrose solutions by osmotic distillation’, J Membrane Sci, 170, 281–289. Courel M, Dornier M, Rios GM and Reynes M (2000a), ‘Modelling of water transport
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280 Separation, extraction and concentration processes in osmotic distillation using asymmetric membrane’, J Membrane Sci, 173, 107– 122. Courel M, Tronel-Peyroz E, Rios GM, Dornier M and Reynes M (2001), ‘The problem of membrane characterization for the process of osmotic distillation’, Desalination, 140, 15–25. Curcio E, Barbieri G and Drioli E (2000), ‘Operazioni di distillazione a membrane nella concentrazione dei succhi di frutta’, Industria delle Bevande, XXIX, 113–121. Deblay P (1995), ‘Process for at least partial dehydration of an aqueous composition and devices for implementing the process’, US Patent 5,382,365. Diban N, Voinea OC, Urtiaga A and Ortiz I (2009), ‘Vacuum membrane distillation on the main pear aroma compound: experimental study and mass transfer modelling’, J Membrane Sci, 326, 64–75. Ding Z, Ma R and Fane AG (2002), ‘A new model for mass transfer in direct contact membrane distillation’, Desalination, 151, 217–227. Drioli E, Jiao B and Calabrò V (1992), ‘The preliminary study on the concentration of orange juice by membrane distillation’, Proceedings of VII International Citrus Congress, Acireale (Italy), 3, 1140–1144. Duan SH, Ito A and Ohkawa A (2001), ‘Removal of trichloroethylene from water by aeration, pervaporation and membrane distillation’, J Chem Eng Jpn, 34, 1069–1073. Durham RJ and Nguyen MH (1994), ‘Hydrophobic membrane evaluation and cleaning for osmotic distillation of tomato puree’, J Membrane Sci, 87, 181–189. El-Bourawi MS, Ding Z, Ma R and Khayet M (2006), ‘A framework for better understanding membrane distillation separation process’, J Membrane Sci, 285, 4–29. Gabelman A and Hwang S (1999), ‘Hollow fiber membrane contactors’, J Membrane Sci, 159, 61–106. Gabino F, Belleville MP, Preziosi-Belloy L, Dornier M and Sanchez J (2007), ‘Evaluation of the clearing of a new hydrophobic membrane for osmotic evaporation’, Sep Purif Technol, 55, 191–197. Galaverna G, Di Silvestro G, Cassano A, Sforza S, Dossena A, Drioli E and Marchelli R (2008), ‘A new integrated membrane process for the production of concentrated blood orange juice: effect on bioactive compounds and antioxidant activity’, Food Chem, 106, 1021–1030. Garcia-Payo MC, Rivier CA, Marison IW and Stockar UV (2002), ‘Separation of binary mixtures by thermostatic sweeping gas membrane distillation: II. Experimental results with aqueous formic acid solutions’, J Membrane Sci, 198, 197–210. Geankoplis CJ (1993), ‘Principles of mass transfer’, in Transport processes and unit operation, London, Prentice-Hall, 381–413. Godino MP, Peña L, Ortiz de Zárate JM and Mengual JI (1995), ‘Coupled phenomena membrane distillation and osmotic distillation through a porous hydrophobic membrane’, Sep Sci Technol, 30, 993–1011. Gostoli C (1999), ‘Thermal effects in osmotic distillation’, J Membrane Sci, 163, 75–91. Gryta M (2002), ‘Concentration of NaCl solution by membrane distillation integrated with crystallization’, Sep Sci Technol, 37, 3535–3558. Gryta M (2005), ‘Osmotic MD and other membrane distillation variants’, J Membrane Sci, 246, 145–156. Gunko S, Verbych S, Bryk M and Hilal N (2006), ‘Concentration of apple juice using direct contact membrane distillation’, Desalination, 190, 117–124. Herron JR, Beaudry EG, Jochums CE and Medina LE (1994), ‘Osmotic concentration apparatus and method for direct osmotic concentration of fruit juice’, US Patent 5,281,430. Ho WSW and Sirkar KK (1992), Membrane handbook, Chapman & Hall, New York. Hogan PA, Canning RP, Peterson PA, Johnson RA and Michaels AS (1998), ‘A new option: osmotic distillation’, Chem Eng Prog, 94, 49–61.
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Advances in membrane-based concentration 281 Hongvaleerat C, Cabral LMC, Dornier M, Reynes M and Nigsanond S (2008), ‘Concentration of pineapple juice by osmotic evaporation’, J Food Eng, 88, 548–552. Jariel O, Reynes M, Courel M, Durand N, Dornier M and Deblay P (1996), ‘Comparison of some fruit juice concentration techniques’, Fruits, 51, 437–450. Jiao B, Cassano A and Drioli E (2004), ‘Recent advances on membrane processes for the concentration of fruit juices: a review’, J Food Eng, 63, 303–324. Jiao B, Molinari R, Calabrò V and Drioli E (1991), ‘Application of membrane operations in concentrated citru juice processing’, Agro-Ind Hi-Tech, 19–27. Karode SK, Kulkarni SS and Ghorapade MS (2000), ‘Osmotic dehydration coupled reverse osmosis concentration: steady-state model and assessment’, J Membrane Sci, 164, 277–288. Khayet M, Godino MP and Mengual JI (2000), ‘Theory and experiments on sweeping gas membrane distillation’, J Membrane Sci, 165, 261–272. Khayet M, Mengual JI and Matsura T (2005), ‘Porous hydrophobic/hydrophilic composite membranes. Application in desalination using direct contact membrane distillation’, J Membrane Sci, 252, 101–113. Khayet M, Mengual JI and Zakrewska-Trznadel G (2003), ‘Theoretical and experimental studies on desalination using the sweeping gas membrane distillation’, Desalination, 157, 297–305. Khayet M, Mengual JI and Zakrewska-Trznadel G (2006), ‘Direct contact membrane distillation for nuclear desalination. Part II. Experiments with radioactive solutions’, Int J Nuclear Desalination, 56, 56–73. Koroknai B, Gubicza L and Bélafi-Bakó K (2006a), ‘Coupled membrane process applied to fruit juice concentration’, Chem Pap, 60, 399–403. Koroknai B, Kiss K, Gubicza L and Bélafi-Bakó K (2006b), ‘Coupled operation of membrane distillation and osmotic evaporation in fruit juice concentration’, Desalination, 200, 526–527. Kozák A, Bánvölgyi S, Vincze I, Kiss I, Békássy-Molnár E and Vatai G (2008), ‘Comparison of integrated large scale and laboratory scale membrane processes for the production of black currant juice concentrate’, Chem Eng Proc, 47, 1171–1177. Kozák A, Békássy-Molnár E and Vatai G (2009), ‘Production of black-currant juice concentrate by using membrane distillation’, Desalination, 241, 309–314. Laganà F, Barbieri G and Drioli E (2000), ‘Direct contact membrane distillation: modelling and concentration experiments’, J Membrane Sci, 166, 1–11. Lawson KW and Lloyd DR (1997), ‘Membrane distillation’, J Membrane Sci, 124, 1–25. Lebfevre MSM (1988), ‘Method of performing osmotic distillation’, US Patent 4,781,837. Luh BS, Feinberg B, Chung JI and Woodroof JG (1986), ‘Freezing fruits’ in Woodroof JG and Luh BS, Commercial fruit processing, Westport, AVI Publishing Co., 263–351. Maccarone E, Campisi S, Cataldi Lupo MC, Fallico B and Nicolosi Asmundo C (1996), ‘Thermal treatments effects on the red orange juice constituents’, Ind Bevande, 25, 335–341. Mansouri J and Fane AG (1999), ‘Osmotic distillation of oily feeds’, J Membrane Sci, 153, 103–120. Mengual JI, Ortiz de Zárate JM, Peña L and Velázquez A (1993), ‘Osmotic distillation through macroporous hydrophobic membranes’, J Membrane Sci, 82, 129–140. Mengual JI and Peña L (1997), ‘Membrane distillation’, Colloid Interf Sci, 1, 17–29. Michaels AS (1998), ‘Methods and apparatus for osmotic distillation’, US Patent 5,824,223. Michaels AS (1999), ‘Osmotic distillation process using a membrane laminate’, US Patent 5,938,928. Milleville H (1990), ‘Direct osmotic concentrates juices at low temperature’, Food Proc, 51, 70–71. © Woodhead Publishing Limited, 2010
282 Separation, extraction and concentration processes Mulder M (1998), Basic principles of membrane technology, 2nd edn. London, Kluwer Academic Publishers, 280–303. Nagaraj N, Patil G, Babu BR, Hebbar UH, Raghavarao KSMS and Nene S (2006), ‘Mass transfer in osmotic membrane distillation’, J Membrane Sci, 268, 48–56. Narayan AV, Nagaraj N, Hebbar HU, Chakkaravarthi A and Raghavarao KSMS (2002), ‘Acoustic field-assisted osmotic membrane distillation’, Desalination, 147, 149–156. Ortiz de Zárate JM, Rincón C and Mengual JI (1998), ‘Concentration of bovine serum albumin aqueous solutions by membrane distillation’, Sep Sci Technol, 33, 283–296. Petrotos KB, Quantick PC and Petropakis H (1998), ‘A study of the direct osmotic concentration of tomato juice in tubular membrane module configuration. I. The effect of certain basic process parameters on the process performance’, J Membrane Sci, 150, 99–110. Petrotos KB, Quantick PC and Petropakis H (1999), ‘Direct osmotic concentration of tomato juice in tubular membrane module configuration. II. The effect of using clarified tomato juice on the process performance’, J Membrane Sci, 160, 171–177. Petrotos KB and Lazarides HN (2001), ‘Osmotic concentration of liquid foods’, J Food Eng, 49, 201–206. Ravindra Babu B, Rastogi NK and Raghavarao KSMS (2006), ‘Mass transfer in osmotic membrane distillation of phycocyanin colorant and sweet-lime juice’, J Membrane Sci, 272, 58–69. Ravindra Babu B, Rastogi NK and Raghavarao KSMS (2008), ‘Concentration and temperature polarization effects during osmotic distillation’, J Membrane Sci, 322, 146–153. Rektor A, Vatai G and Békássy-Molnár E (2006), ‘Multi-step membrane processes for the concentration of grape juice’, Desalination, 191, 446–453. Rodrigues RB, Menezes HC, Cabral LMC, Dornier M, Rios GM and Reynes M (2004), ‘Evaluation of reverse osmosis and osmotic evaporation to concentrate camu-camu juice (Myrciaria dubia)’, J Food Eng, 63, 97–102. Romero J, Draga H, Belleville MP, Sanchez J, Come-James C, Dornier M and Rios GM (2006), ‘New hydrophobic membranes for contactor processes – applications to isothermal concentration of solutions’, Desalination, 193, 280–285. Sakai K, Muroi T, Ozawa K, Takesawa S, Tamura M and Nakaue T (1986), ‘Extraction of solute-free water from blood by membrane distillation’, Trans Am Soc Artif Intern Organs, 32, 397–400. Schneider K, Holz W and Wollbeck R (1988), ‘Membranes and modules for transmembrane distillation’, J Membrane Sci, 39, 25–42. Schofield RW, Fane AG and Fell CJD (1987), ‘Heat and mass transfer in membrane distillation’, J Membrane Sci, 33, 299–313. Shaw PE, Lebrun M, Dornier M, Ducamp MN, Courel M and Reynes M (2001), ‘Evaluation of concentrated orange and passion fruit juices prepared by osmotic evaporation’, Lebensm Wiss Technol, 34, 60–65. Sheng J, Johnson RA and Lefebvre MS (1991), ‘Mass and heat transfer mechanism in the osmotic distillation process’, Desalination, 80, 113–121. Sirkar KK (1997), ‘Membrane separation technologies: current developments’, Chem Eng Commun, 157, 145–184. Thompson D (1991), ‘The application of osmotic distillation for the wine industry’, Aust Grapegrower Winemaker, 11, 11–14. Tomaszewska M (2000a), ‘Concentration and purification of fluosilicic acid by membrane distillation’, Ind Eng Chem Res, 39, 3028–3041. Tomaszewska M (2000b), ‘Membrane distillation – examples of applications in technology and environmental protection’, Pol J Environ Stud, 9, 27–36. Tomaszewska M, Gryta M and Morawski AW (1995), ‘Study on the concentration of acids by membrane distillation’, J Membrane Sci, 78, 277–282. © Woodhead Publishing Limited, 2010
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284 Separation, extraction and concentration processes
10 Separation of value-added bioproducts by colloidal gas aphrons (CGA) flotation and applications in the recovery of value-added food products P. Jauregi and M. Dermiki, The University of Reading, UK Abstract: This chapter describes the application of a flotation method based on the formation of colloidal gas aphrons (CGA), which are surfactant-based microbubbles, to the recovery of value-added food products from fermentation broths and waste streams of food and agricultural industries. Their main properties are described and contrasted with those of conventional foams. The fundamentals of the separation are described and illustrated by several case studies on CGA applications to the recovery of whey proteins, polyphenols and carotenoids. Finally, a critical evaluation of the feasibility of CGA for industrial applications is carried out based on the astaxanthin case study and future applications of CGA are discussed. Key words: colloidal gas aphrons, flotation, bioproducts, surfactants, food ingredients, recovery.
10.1 Introduction Many studies have been conducted in recent years into additional biological functionalities of several food components and ingredients such as the dairy proteins, peptides, polyphenols or carotenoids so that they could be produced as ingredients in functional foods. The development of functional foods poses new challenges to the food industry as new processes and technologies are required. The food industry must adapt to these changing times and to the revolution caused by newly emerging functional foods. In addition, in the last few years reducing the environmental impact of industrial wastes has been a subject of increasing concern. In industrial wastewaters, these compounds considerably increase biochemical and oxygen demands, with detrimental effects on the flora and fauna of discharge zones, whereas in solid residues © Woodhead Publishing Limited, 2010
Separation of value-added bioproducts by colloidal gas aphrons 285 used as fertiliser they may inhibit germination. Furthermore sustainability is given high priority by governments and industries are encouraged to develop practices for the usage and recycling of wastes. Food processing waste streams and agricultural wastes can be exploited as cheap sources of high-value products. For example, proteins and peptides can be extracted from whey which is a by-product in cheese production and polyphenols can be extracted from winemaking waste, such as grape skin and marc. Separation and purification of these products is carried out by a number of steps including liquid–liquid extraction, membrane techniques and/or chromatographic methods. Alternative separations such as flotation may be used for the recovery stage although further processing will be necessary if high purity products are required. Foam fractionation has been applied for the recovery of proteins. Flotation, which is the foam separation of insoluble compounds, has been applied mainly to wastewater treatment and to the recovery of minerals. Separation of biomass has also been successfully carried out by flotation. This chapter deals mainly with a particular type of flotation using microbubbles which are also called colloidal gas aphrons (CGA). CGA were first described by Sebba (1972) as surfactant-stabilised microbubbles (10–100 mm) generated by intense stirring (>8000 rpm) of a surfactant solution. In this chapter, the application of CGA to bioseparations is reviewed particularly in relation to the recovery of value-added food products. First, the structural and dispersion characteristics of CGA are described. Then, the separation fundamentals are described and illustrated with the application of CGA to the separation of proteins. In addition, the feasibility of CGA for industrial applications is considered, focusing on: (i) scalability using a flotation column and (ii) removal and recycling of surfactant. All these aspects are illustrated in several case studies of CGA separations for the recovery of bioactives from plant extracts (norbixin and polyphenols) and recovery of astaxanthin produced by microbial fermentation.
10.2 Colloidal gas aphrons (CGA) properties 10.2.1 Structure of CGA Colloidal gas aphrons (CGA) are surfactant-stabilised microbubbles (10– 100 mm) generated by intense stirring of a surfactant solution at high speeds (>8000 rpm). Sebba (1987) postulated that they are composed of multilayers of surfactant molecules; as depicted in Fig. 10.1 surfactant molecules adsorb at the interface with hydrophilic heads towards the aqueous phase and the hydrophobic tails towards the gas phase. Sebba’s hypothesis on the structure was based on several experimental observations such as delayed coalescence and hence higher stability of CGA than conventional foams. Jauregi et al. (2000) investigated the drainage kinetics of CGA and compared measured drainage rates with those obtained by applying predictive models for foams with and without modifications. The main modifications were in relation to © Woodhead Publishing Limited, 2010
286 Separation, extraction and concentration processes Inner surface of shell
Air core
Viscous water Shell
Electrical double layer
Outer surface of shell
Fig. 10.1 Proposed structure of CGA by Sebba (1987).
structural differences between foams and CGA upon drainage, such as aphrons being surrounded by a liquid film so that they do not adopt a dodecahedral shape. Interestingly, the model giving the best prediction was the one including differences in structural features which further supports Sebbas’s theory. Jauregi and coworkers used for the first time small angle x-ray diffraction in an attempt to determine the thickness of the surfactant film and the number of surfactant layers of CGA generated by the anionic surfactant sodium bis(2-ethyl hexyl) sulfosuccinate (AOT). The analysis of the data proved to be difficult as for the same sample different film thickness values could be obtained. However, an interesting finding was that samples containing aphrons gave similar scattering regardless of the surfactant concentration and these corresponded to 5.4 nm (which is equivalent to seven layers of surfactant assuming the full length of the surfactant molecule arranged in layers and vertically at the interface). Moreover, the same surfactant solutions with no aphrons gave a different scattering signal and, in this instance, differences between different concentrations of surfactant solutions were found; for example, at concentrations around the cmc (critical micellar concentration), the scattering corresponded to a bilayer, which could correspond to micelles whereas, above the cmc, the scattering corresponded to a multilayer of 3–5 molecules, confirming the hypothesis that AOT forms lamellar structures (Jauregi et al., 2000). In a number of research studies predictive models for liquid drainage in CGA were developed based on measurements of liquid drainage rate (Table 10.1). Most of the studies highlight the difference between CGA and conventional foams. Bhatia et al. (2005) studied the effect of stirring speed and stirring time on the stability of CGA generated using a range of
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Table 10.1 Summary of studies for the characterisation of CGA (c, cationic; a, anionic; n-i, nonionic) (adapted from Dermiki, 2009) Surfactant
Parameters studied
Main findings
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Zhao et al., SDS (a) 2009; Larmignat CTAB (c) et al., 2008 Tween 20 (n-i)
Surfactant type, Csurf, pressure CGA rheology is not affected by pipe shape or hydraulic diameter, CGA drop, T, volumetric and mass can be treated as shear thinning fluid for all surfactant solutions. Increase flow rate, pipe diameters and in Csurf leads to increased shear stress for a given shear rate shape
Feng et al., 2009 Tergitol (n-i) Ramnolipid (a)
Csurf, pH, Csalt
Stability increases with Csurf and decreases with increasing pH, Csalt. Three distinct phases in the liquid drainage
Moshkelani and SDS (a) Amiri, 2008
Electrical conductivity of surfactant solution
Three clear stages of drainage of CGA can be identified, compared with two stages for conventional foams
Tseng et al., 2006
Pressure drop, T, mass and volumetric flow rate of CGA
CGA can be considered shear thinning fluid. Heat transfer coefficients for CGA made of water and surfactant are smaller than that of water, and they are constant and independent of mass flow and heat flux
tstir, rpm, Csurf, mixing two oppositely charged CGA
Development of an empirical correlation for variation of gas hold-up with tstir, rpm Ê t ˆ – 4.025ˆ a = (0.004 + 44.017) Á 0.2407 ln ÊÁ –2.00 ˜¯ ˜¯ Ë2 Ë
Bhatia et al., 2005
SLS (a) CTAB (c)
The mixing of the two oppositely charged CGA showed no effect on drainage Yan et al., 2005 SDS (a) CTAB (c) Tween 80 (n-i)
Type of surfactant, Csurf, T
There was no effect of the type of surfactant on the stability. Mathematical rn model which describes drainage of CGA: Vt = Vmax n . Two distinct K + rn stages of drainage are determined by two independent mechanisms
Jarudilokkul et al., 2004
Tween 20 (n-i)
tstir, pH, rpm, CNaCl, Csurf, protein separation
Increasing Csurf, tstir increases the stability Increase of Csurf leads to decrease of protein separation
Dai and Deng, 2003
HTAC (c)
pH, concentration of silica sol for stabilisation
CGA were stable for up to 12 h at concentration of silica sol 0.15–0.25 mol dm–3 and pH 7–10
Separation of value-added bioproducts by colloidal gas aphrons 287
Reference
© Woodhead Publishing Limited, 2010
Reference
Surfactant
Parameters studied
Main findings
Jauregi et al., 2000
AOT (a)
Csurf, pH, Csalt, T, tstir
X-ray diffraction for the characterisation of CGA showed evidence of the existence of surfactant multilayers
Jauregi et al., 1997
AOT (a)
Csurf, pH, salt, T, tstir
Increase of Csurf increases the stability, whereas addition of salt leads to the opposite effect Significant interactive effects: (i) Csurf · salt, (ii) pH · Csurf, (iii)T · tstir, (iv) Csurf · T Gas hold-up depends on tstir, Csurf, and salt
Bredwell et al., 1995
SDS (a) Csurf, salt CPC (c) Triton X-100 (n-i)
Save and DTAC, CTAC, Pangarkar, 1994 CPC, DMDSAC (c) SDBS, SLS (a) Chaphalkar et al., 1993
pH, tstir, Csurf, Csalt, viscosity, additives, impeller clearance
SDBS (a), CTAB Type and concentration of (c), surfactant, ionic strength Tergitol (n-i)
Amiri and TTAB (c) Woodburn, 1990
Csurf, pH
Increasing Csurf leads to decrease of formation time of CGA and increase in CGA stability. No effect of salt. No effect of surfactant type and surfactant concentration on the gas hold-up pH and impeller clearance had no effect on stability. Type of surfactant, viscosity, addition of enzymes, polymers, nonionic surfactants, solvents and salts affect gas hold-up and stability MDtergitol<MDCTAB<MDDDBS. Increase of Csurf leads to decrease of MD. Addition of salts reduces the MD of ionic surfactants There is agreement between the predicted and observed rise velocities Stability of CGA depends on the Csurf and pH
MD, mean diameter; tstir, stirring time; Csurf, concentration of surfactant; Csalt, concentration of salt; T, temperature; rpm, stirring speed. SDS, sodium dodecyl sulfate; SDBS, sodium dodecylbenzenesulfonate; CTAB, cetyltrimethyl ammonium bromide; SLS, sodium lauryl sulfate; AOT, sodium bis (2-ethyl hexyl) sulfosuccinate. HTAC, hecadecyl trimethyl ammonium chloride; CPC, cetyl piridinium chloride; DTAC, dodecyltrimethylammonium chloride; CTAC, cetyltrimethylammonium chloride; DMDSAC, dimethyl distearyl ammonium chloride; TTAB, tetradecyl trimethyl ammonium bromide.
288 Separation, extraction and concentration processes
Table 10.1 Continued
Separation of value-added bioproducts by colloidal gas aphrons 289 surfactants and they developed a first order kinetics model for the drainage. Interestingly, drainage kinetics did not change when they used a mixture of an anionic and a cationic surfactant as opposed to conventional foams (Bhatia et al,. 2005). Yan et al (2005) developed a mathematical kinetic model by studying CGA generated using a range of concentrations of surfactants at varying temperatures. According to this model there were two distinct stages of CGA drainage (Yan et al., 2005). On the other hand, in a more recent study Moshkelani and Amiri (2008) measured the electrical conductivity of CGA dispersions and found that there are three separate stages in the liquid drainage of aphrons as opposed to the two stages in conventional foams (Moshkelani and Amiri, 2008). This was further supported in a more recent study by Feng et al. (2009) where the volume of drained liquid was measured and also photographs of the foams at different stages were taken. At the first stage they showed how drainage increases with time owing to a combination of upflow migration of bubbles and downward liquid drainage under gravity. Then, in the following stage, the dispersion behaves like conventional foams and, at the final stage, the foam becomes water deficient and begins to behave like dry foam. At this stage, the foam drainage is very slow owing to slow liquid release from films under capillary suction , i.e.: liquid is released from the film owing to lower pressure in the capillaries between the bubbles (Feng et al., 2009). Overall these studies support further the hypothesis that there are structural differences between CGA and conventional foams. 10.2.2 Characteristics of CGA Owing to their unique structure CGA possess the following important properties: ∑
∑
Higher stability than conventional foams. Owing to the multilayer structure, CGA exhibit high stability. When two aphrons collide, the momentum may not be enough to break the barrier of six surfactantstabilised interfaces, as opposed to the two surfactant-interfaces when two bubbles of conventional foams collide (Sebba, 1987). Stability is measured in terms of the liquid drainage rate, which is generally considered to follow first order kinetics (Bhatia et al., 2005) consequently, half-life (t) is defined as the time required for half the original volume of liquid to drain. CGA are also characterised in terms of gas hold-up (e), which is defined as the ratio between the gas volume (Vg) and the dispersion final volume after stopping stirring and at time = 0 of drainage (Va0): e = Vg/Va0
[10.1]
The buoyancy of the encapsulated gas leads to easy separation of the aphron phase from the bulk liquid phase. Therefore, no centrifuges are needed to separate the two phases. On the other hand, creaming can be
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290 Separation, extraction and concentration processes avoided when necessary by stirring CGA at such a rate that the lateral movement conveyed to the bubbles is greater than the upward buoyancy owing to gravity. ∑ Adherence of particles and molecules to the encapsulating shell. It is possible to modify the surface properties of CGA by using different types of surfactant. Depending on the surfactant used to generate the CGA, the outer surface of the bubble can be negatively (anionic surfactant), positively (cationic surfactant) or non-charged (nonionic surfactant). Consequently, oppositely charged molecules adsorb. Thus, the selectivity of adsorption can be modified by changing the type of surfactant. This is illustrated in experiments carried out in our laboratory with two oppositely charged dyes, the cationic methylene blue, and the anionic methyl orange and CGA generated by the cationic surfactant CTAB (see Plate I, between pages 292 and 293). CGA are contacted with the aqueous solution containing both dyes and then the dispersion is allowed to settle so it separates into two phases (Plate Ic). The top CGA phase contains mainly methyl orange (orange colour top phase) and the liquid phase contains mainly methylene blue (blue colour liquid phase). Basu and Malpani (2001) also reported the effective separation of these dyes using CGA generated by CTAB. ∑ Low viscosity of the system. Flow properties are similar to those of water, as long as the gas hold-up does not exceed 65% (Roy et al., 1995), regardless of the type of the surfactant used to generate them. Consequently, CGA can be pumped easily from the generation point to the point where they are going to be used, whereas, for conventional foams, their characteristics change during pumping owing to the elastic nature of the bubble. This is advantageous when applied to a flotation column for fractionation and flotation in the removal and/or recovery of products. Several researchers have investigated the rheology of CGA, as shown in Table 10.1 (Larmignat et al., 2008; Tseng et al., 2006; Zhao et al., 2009). An interesting application which takes advantage of CGA rheological properties is described in a patent by Brookey (2004) where CGA are used to produce fluids with improved shear thinning properties, and this has led to a successful industrial application for well-drilling fluids by MASI technologies LLC. ∑ Small size of bubbles, resulting in larger interfacial area per unit volume, and thus a large capacity for adsorption of molecules and fine particles. Bubble size depends on the concentration and type of surfactant, ionic strength and the presence of other molecules or particles. Jauregi (1997) found that bubble size of AOT-generated CGA increased with surfactant concentration, ionic strength and stirring time up to 10 min. Size distribution of CGA generated by selected surfactants can be seen in Table 10.2.
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Separation of value-added bioproducts by colloidal gas aphrons 291 Table 10.2 Characteristics of CGA generated by ionic and nonionic surfactants (values found at optimum conditions) Surfactant
cmc (mM)
Gas hold-up
Half life (s)
CTAB (c)
0.9 (in H2O)
0.70 <500 40–260 (Fuda et al., 2005) (Fuda et al., 2005; (Chaphalkar Save and Pangarkar, et al., 1993) 1994; Yan et al., 2005)
Tween 20 (n-i) 0.056
0.4–0.63 (Dermiki, 2009)
100–600 (Dermiki, 2009)
Tween 60 (n-i) 0.027
0.10–0.60 (Dermiki et al., 2009)
100–500 (Dermiki et al., 2009)
Bubble size (mm)
AOT (a)
0.84 0.02–0.65 (Fuda, 2004) (Jauregi, 1997)
30–930 46–101 (Jauregi et al., 1997) (Jauregi et al., 2000)
SDS (a)
8.08
<550 (Yan et al., 2005)
0.70 (Matsuhita et al., 1992)
10.2.3 Generation of CGA Formation of aphrons The formation of CGA requires, as described by Sebba (1985), a horizontal disc that rotates at very high speeds. Baffles are also recommended in order to achieve the required mixing regime and produce smaller bubbles (Jauregi et al., 1997). Furthermore, Jauregi et al. (1997) determined the power input required to generate CGA and found that there is a minimum power requirement to generate CGA, which is dependent on the surfactant concentration as this affects the volume of air incorporated; for surfactant concentrations below and around the cmc the minimum power requirement to generate CGA was 45 kW m–3. In our laboratory, CGA are typically generated by stirring a surfactant solution with a laboratory mixer at 8000 rpm. A four-bladed impeller with a high shear screen (Fig. 10.2) was found to generate CGA dispersions with small bubble sizes (average diameters ranged between 35 and 70 mm) (Jauregi et al., 1997). Some other apparatus reported by Sebba (1985) involved the use of a venturi throat at which gas is admitted but this is not satisfactory for production of CGA on a large scale. In a recent study by Xu et al. (2008), the use of sonication for the formation of CGA was investigated and compared with mechanical agitation, which is the most widely used method. They found that sonication led to CGA with higher gas hold-up, smaller bubble size, higher number of bubbles and larger interfacial area than mechanical agitation; processing time was also shorter. However, sonication is an expensive method that cannot be easily scaled up.
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292 Separation, extraction and concentration processes
Fig. 10.2 Impeller used with laboratory mixer for the generation of CGA.
Type of surfactants used in the generation of CGA A wide range of surfactants can be used for the generation of CGA: anionic, e.g. sodium dodecyl sulfate (SDS); cationic, e.g. cetyltrimethylammonium bromide (CTAB); nonionic, e.g. Tween 20. Characteristics of CGA generated with these surfactants are summarised in Table 10.2; also an overview of the characteristics of CGA generated from different surfactant solutions is given by Jauregi et al. (1997). The use of surfactants for food applications can be problematic particularly if they are as toxic as the ionic surfactants. Nonionic surfactants on the other hand can be advantageous to use as they are non-toxic and they are used in the formulation of medicines; hence, they may not need to be removed from the CGA phase and at the same time could help to formulate the final product (section 10.3.1). The use of natural surfactants such as saponin extracted from plants has also been explored (Kommalapati et al., 1998). The biosurfactant rhamnolipid produced by Pseudomonas aeroginosa (Feng et al., 2009) is an alternative to the synthetic ones because biosurfactants are biodegradable and hence have a less harmful impact on the environment. However, these surfactants, although harmless to the environment, can, in high concentrations, be toxic so their use in food applications needs to be further studied. Moreover, not all surfactants are suitable for generation of CGA. For instance, those surfactants with poor
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Separation of value-added bioproducts by colloidal gas aphrons 293 foaming ability cannot produce CGA. Surfactants that have a tendency to form very stable micelles or vesicles have poor foaming ability. Micellar stability is inversely related to foaming ability because very stable micelles are less capable of providing the flux of surfactants necessary to stabilise the new air–solution interface created during foaming (Rosen, 2004). In experiments carried out in our laboratory with the less toxic cationic surfactant dioctadecyldimethylammonium bromide (DODAB) at 0.1 and 1 mM aqueous solutions, we could not generate CGA. This surfactant has the ability to form very stable vesicles owing to its structure and, hence, has very low foaming ability. However, it would be very useful to generate CGA using DODAB because it is used in medicinal applications and may be suitable for food applications. Span surfactants have a low foaming ability whereas their nonionic ethoxylated derivatives, Tween surfactants, have a relatively high foaming ability. CGA formed in combination with those two surfactants led to decreased CGA stability compared with the ones produced only from the Tween surfactants (Dermiki et al., 2009) (section 10.3.2, Table 10.4). Worden and Scranton (2000) describe an interesting method to produce reversible gas or liquid aphrons using a combination of polymers as surfactants such as methacrylic acid and polyethylene glycol. The method is based on the polymers changing their foaming and emulsifying properties with pH thus a fast coalescence of CGA is achieved by changing the pH. This can be of particular interest for applications such as waste-treatment and for increasing mass transfer in bioreactors. However, for bioseparations the removal of surfactant still remains a problem.
10.3 Applications of CGA in the recovery of value-added food products As a consequence of the above properties, researchers have considered various applications for CGA, with a particular focus on separation processes (Table 10.3). For example, CGA have been used for the flotation of biological products such as microbial cells and proteins; see also review by Jauregi and Varley (1999) on applications of CGA in biotechnology. Other interesting biotechnological applications include the use of CGA in bioprocesses to enhance gas mass transfer in fermentations (Bredwell and Worden, 1998; Kaster et al., 1990; Park et al., 2009) or even in wastewater treatment (Dai et al., 2004). A number of recent research studies focus as well on the environmental applications of CGA such as the remediation of soil (Boonamnuayvitaya et al., 2009; Chu, 2003; Couto et al., 2009; Kommalapati et al., 1998; Roy et al., 1995) and the removal of oil and dyes from wastestreams (Roy et al., 1992) as summarised in Table 10.3. High recoveries of proteins and other bioproducts indicate the potential of CGA in downstream processing, where alternative methods of reducing
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Table 10.3 Summary of applications of CGA in separation processes (adapted from Dermiki et al., 2008) Surfactant
Principles/driving forces of the separation
References
Recovery of antioxidants: astaxanthin
CTAB (c)
Recovery is driven by electrostatic interactions
(Dermiki et al., 2008) (Dermiki et al., 2009) (Alves et al., 2006) (Spigno and Jauregi, 2005)
Highest recovery and enrichment factor under conditions that favour electrostatic interactions. Important effect of flocculating agent
(Zidehsaraei et al., 2009) (O’Connell and Varley, 2001)
Recovery is driven by electrostatic interactions Electrostatic interactions are important for the recovery, the selectivity can be manipulated by changing conformation of proteins Continuous separation more efficient than batch Hydrophobic interactions play an important role
(Fuda et al., 2005) (Fuda et al., 2004) (Fuda and Jauregi, 2006)
Better recovery with the ionic surfactants because they produce more stable dispersions
(Mansur et al., 2006)
Recovery is driven by mechanism of bubble flotation
(Mansur et al., 2004)
The adsorption of yeast on CGA follows the Langmuir model. Changes in pH and feed concentration lead to a change of mechanism of bubble attachment and detachment from a monolayer to a multilayer adsorption. The volume of air incorporated in the system is a significant factor for separation
(Hashim et al., 1998) (Hashim et al., 1995) (Hashim et al., 2000)
norbixin gallic acid Recovery of enzymes: glucoamylase lipase immobilisation
CTAB (c) CTAB (c) TTAB (c) SDS (a)
Recovery of proteins: CTAB (c) b-lactoglobulin lactoferrin, lactoperoxidase AOT (a) CTAB (c), AOT (a) whey proteins whey proteins lysozyme, b-casein Removal of fine particles
Clarification of yeast cells
SLS (a) TWEEN 20, TWEEN 40, TWEEN 60, TWEEN 80 (n-i) HTAB (c) SDBS (a) TWEEN 20 (n-i) HTAB (c) SDBS (a) TWEEN 20 (n-i) BDHA (c)
(Amiri and Valsaraj, 2004) (Jarudilokkul et al., 2004)
294 Separation, extraction and concentration processes
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Application
Clarification of suspensions: oil mill effluent suspension of microalgae suspension of inorganic minerals
AOT (a), SDBS (c), LUX flakes (a), BDHA (c)
Recovery is enhanced when electrostatic interactions are (Basu and Malpani, 2001) favoured (Roy et al., 1992) The height of the separation column is an important factor in (Hashim et al., 1999) order to ensure the required separation CGA better oil removal efficiencies. Removal of oil positively affected by initial oil concentration and particle diameter CGA more efficient than conventional surfactant solution CGA more efficient than conventional surfactant solutions Nonionic surfactant better for the removal of naphthalene because it has better solubilising power for the compound CGA more efficient than simple waterflood, the recovery results from solubilisation of the hydrophobic compound in the surfactant Maximum effectiveness of separation was achieved at pH close to the pK of each surfactant, removal effectiveness increased when CGA were used at the temperature at which they were produced, rather that at high temperatures, the effectiveness of solids removal was a function of air-tosolids ratio and of the solids concentration
(Couto et al., 2009) (Boonammuayvitaya et al., 2009) (Roy et al., 1994) (Roy et al., 1995) (Kommalapati et al., 1998) (Subramaniam et al., 1990)
SDS, sodium dodecyl sulfate; SDBS, sodium dodecylbenzenesulfonate; SLS, sodium lauryl sulfate; AOT, sodium bis (2-ethyl hexyl) sulfosuccinate; HTAB (= CTAB), hexadecyltrimethylammonium bromide; TTAB, tetradecyltrimethylammonium bromide; BDHA, benzyldimethylhexadecylammonium chloride.
Separation of value-added bioproducts by colloidal gas aphrons 295
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Remediation of wastewater: removal of organic dyes HTAB (c), SDS (a), SDBS (a) clarification of oily SDS (a) wastewater Remediation of soil: remediation of sandy soil SDS (a) contaminated with diesel oil BioNonex and Biosolve removal of pyrene from soil SDS (a) hazardous oily waste Tergitol (n-i), HTAB flushing naphthalene (c), SDS (a) Plant-based surfactant removal of hexachlorobenzene (HCB)
296 Separation, extraction and concentration processes processing costs are continuously sought. This is particularly important in the extraction of value-added products from food and agricultural waste. Research carried out by our group shows that CGA can successfully recover proteins from whey (Fuda et al., 2005; 2004), polyphenols from wine waste extracts (Spigno and Jauregi, 2005; Spigno et al., 2010) and carotenoids such as norbixin from plant extracts (Alves et al., 2006). We have also obtained high recovery of astaxanthin particles from a fermentation mixture containing yeast cells (Dermiki et al., 2008). A good understanding of the mechanism of separation is required in order to predict and optimise separations. Therefore, we investigated the mechanism of separation of proteins with CGA generated from ionic surfactants and the mechanism of separation (flotation) of astaxanthin particles, both of which are described in the following subsections. 10.3.1 Recovery of soluble compounds and mechanism of separation Recovery of proteins Similarly to foam fractionation, CGA can be applied for the recovery of soluble compounds. For foam fractionation, the selectivity of separation relies on differences in surface tension of the mixture components whereas in CGA fractionation the selectivity is based on differences in interaction between the surfactant in CGA and the mixture components. Some studies have described the successful separation of proteins with CGA for protein separation. Jauregi and Varley (1999) and Noble and Varley (1999) detailed laboratory scale recovery of proteins from single-protein model solutions (Fig. 10.3). A few studies (Amiri and Valsaraj, 2004; Fernandes et al., 2002b; Fuda et al., 2005) reported the recovery of proteins from a crude extract partly because of a lack of understanding of the mechanism of separation. Our investigation of the separation of proteins from whey has led to an improved insight into the mechanism of separation and to identification of the main operating parameters. Whey contains many proteins that differ in physicochemical properties such as charge, size and hydrophobicity. In a study with CGA generated with an anionic surfactant AOT and whey (Fuda et al., 2004), we found that selective separation of lactoferrin and lactoperoxidase could be achieved at conditions which promote strong electrostatic interactions between the proteins and surfactant (pH= 4). Interestingly, high ionic strength led to higher purity as, at these conditions, drainage of CGA is favoured, resulting in a higher number of contaminant proteins partitioning into the liquid phase. Statistical analysis of the data led to predictive models that correlated purity and enrichment ratio (ratio of protein concentration in the aphron phase to liquid phase) with the main operating parameters which were pH, ionic strength and volume of whey; these models could explain 70% and 62% of the data for purity and enrichment ratios, respectively. Volume of whey is a measure of total protein load and it was found that an increase in volume of whey resulted
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Separation of value-added bioproducts by colloidal gas aphrons 297 Feed suspension
Aphron phase
Liquid phase CGA generator
Fig. 10.3 Schematic representation of the batch small-scale procedure for protein separation using CGA.
in lower separation. This could be partly explained in terms of capacity of CGA but also in terms of protein competition as, at high protein load, an increase in protein–protein interaction can reduce selectivity of separation. A study of separation with model mixtures of lactoferrin–lactoperoxidase (Lf–Lp) and b-lactoglobulin (b-Lg) (Fuda and Jauregi, 2006) demonstrated that the recovery of Lf–Lp decreases from 80 to 45% when in the presence of b-Lg. Further investigations into the mechanism of the separation with CTAB generated CGA and whey confirmed that CGA when generated with ionic surfactants can act similarly to ion exchangers and hence selectivity can be manipulated by the type of surfactant (cationic or anionic), pH and ionic strength of the solution. Thus, under optimum conditions and with CGA, generated by CTAB, selective separation of b-Lg from whey was obtained with 90% b-Lg recovered in the CGA phase and most of the other proteins including BSA and a-lactalbumin remained in solution at basic pH (Fuda et al., 2005). Under these conditions b-Lg interacted strongly with CTAB mainly by electrostatic interactions which resulted in its precipitation. This maximised drainage, which led to the recovery of b-Lg in the form of an insoluble complex with the surfactant and most of the other proteins in the drained liquid. Although results could be explained based on CGA acting as ion exchangers, the poor recovery of proteins with similar surface charge characteristics as b-Lg led to further investigations. Protein–surfactant interactions were further investigated by measurements of zeta potential and fluorescence and these revealed that conformational features of the protein and particularly in relation to denaturation by surfactant molecules also had an influence on the selectivity of the separation. Some studies have been carried out using nonionic surfactants for the © Woodhead Publishing Limited, 2010
298 Separation, extraction and concentration processes recovery of proteins. Jarudilokkul et al. (2004) showed that hydrophobic interactions are important when nonionic surfactants are used for the recovery of proteins such as, lysozyme and b-casein. Noble et al. (1998) applied CGA generated with Triton X-100 to a range of proteins and found that the highest recoveries (74%) were obtained with the most hydrophobic protein (thaumatin). Similarly, Fernandes et al. (2002a) found that the enzyme with a hydrophobic fusion tag was recovered with a higher yield than the wild type cutinase with CGA generated with the nonionic detergent Triton X-114, thus suggesting that hydrophobic interactions were important in the recovery of these proteins. Foam fractionation has also been applied to the recovery of proteins. In this method, bubbles are generated by sparging air through a sintered glass sparger of a given pore size and the proteins are separated based on differences in their surface activity; those with the highest surface activity adsorb to the air-liquid interface of the rising bubbles and separate from the bulk. Noble and co-workers (1998) recovered 86% casein and 25% lysozyme from a binary mixture. Noel et al. (2002) applied a semi-batch foaming process for the recovery of lactoferrin from milk and obtained 40% recovery and a separation ratio of 17.8. Foam fractionation can be a cost effective way of fractionating bioproducts as no additional chemicals are required, but selectivity depends on differences in foaming properties of the components in the mixture and these cannot be easily manipulated by operating parameters such as pH, as foaming of proteins is mainly dependent on their structural properties. Recovery of bioactive components from plant extracts Plant extracts are an important source of value-added products such as polyphenols. In addition, there is much interest in minimising waste and extracting products from agricultural and food waste. We have investigated the recovery of norbixin with CGA from an alkali extract of annatto seeds (Alves et al., 2006). The carotenoid bixin represents more than 80% of the total carotenoids in annatto seeds and is liposoluble. When extracted in alkali conditions the ester group is hydrolysed and becomes the hydrosoluble form called norbixin. It was found that optimum recoveries were achieved at conditions that favour electrostatic interactions using the cationic surfactant CTAB (94%), whereas low recovery (40%) was obtained with the anionic surfactant AOT. The molar ratio of surfactant to norbixin was found to be an important operating parameter. Interestingly maximum recoveries were obtained at CTAB-norbixin molar ratios of 3–4, whereas at higher molar ratios recovery decreased. It was postulated that CTAB and potassium norbixinate interact electrostatically to form an insoluble complex which leads to its effective separation into the CGA phase. In this process, the removal of surfactant was investigated by acidifying the recovered solution to break up electrostatic interactions between surfactant and norbixin molecules and thus reduce the solubility of norbixin which resulted in its precipitation. Bioactive compounds are also found in waste streams of food processes
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Separation of value-added bioproducts by colloidal gas aphrons 299 such as in winemaking. One example is the polyphenols present in grape marc which is a by-product of wine production. Spigno and Jauregi (2005) have investigated the recovery of gallic acid from its standard suspensions with CGA generated from a cationic surfactant. Gallic acid is a polyphenol which becomes ionised and hence negatively charged at pH > 3. Electrostatic interactions were the main driving force for the recovery as optimum recoveries were achieved with the cationic surfactant CTAB at pH = 6. It was found that this pH was also optimum for the recovery of antioxidant activity because at a basic pH loss of antioxidant activity was observed. Further research into the application of CGA to real extracts and using a flotation column showed that at optimum conditions found for the standard gallic acid, high recoveries of total polyphenols were achieved (Spigno et al., 2010). Selectivity of the separation was low as both anthocyanins (expressed as malic acid equivalents) and polyphenols (expressed as gallic acid equivalents) were equally separated into the CGA phase. This led to the use of nonionic surfactants and resulted in high recoveries of all polyphenols (unpublished results). This is an interesting outcome as the use of nonionic surfactant enables integration of the recovery and formulation steps because these surfactants are used in food formulations, e.g. food emulsions (a list of food-grade surfactants is given by Monsalve-Gonzalez and Ochomogo (2009) and a more exhaustive list is given on the website on food additives listed in section 10.6). 10.3.2 Recovery of particles and mechanism of separation Recovery of particles using CGA CGA have been applied for the recovery of fine particles as seen in Table 10.3 and, as observed in those studies, electrostatic interactions had an important effect in the recovery (Waters et al., 2008) as found for soluble compounds (10.3.1). In addition, bubble–particle interactions which are affected by the particle and bubble sizes play a significant role in the recovery of particles. The type of surfactant is important not only because of the surface charge but also because it determines the bubble size of the CGA. Moreover, the concentration of the surfactant will also determine the bubble size and the surfactant molecules available for the recovery of the particles (Mansur et al. 2006). Furthermore, parameters that affect the stability of the CGA dispersion will affect the selectivity of the separation owing to changes in drainage (Waters et al., 2008). Therefore, one can say that the removal of particles is affected by a range of parameters, and this process becomes more challenging for complex systems where more than one type of particle is present. This is illustrated in the case study of the recovery of astaxanthin. Recovery of astaxanthin and mechanism of separation There is increasing interest in the xanthophyll astaxanthin owing to its antioxidant activity and its potential benefits to human health (Hussein et al.,
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300 Separation, extraction and concentration processes 2006). Specifically, natural astaxanthin is produced mainly by two microorganisms: the yeast Phaffia rhodozyma and the microalgae Haematococcus pluvialis. We have investigated the application of CGA following an integrated approach for the recovery of astaxanthin from suspensions of cells, more specifically from the yeast Phaffia rhodozyma (Dermiki et al., 2008). The mechanism of separation was elucidated using a range of surfactants for the generation of CGA and these were subsequently applied for the recovery of astaxanthin from standard suspensions and suspensions of cells (raw extract) as shown in Table 10.4. A minimum recovery of 40% was achieved under all conditions studied, indicating that the recovery was the result of flotation of the hydrophobic astaxanthin particles. However, when electrostatic interactions were promoted at basic pH conditions, recoveries increased further up to 80%, which indicated that CGA were even more efficient in floating the negatively charged particles than the non-charged hydrophobic particles. Interestingly, when the optimum conditions for the standard suspensions of synthetic astaxanthin were applied to the recovery of natural astaxanthin from the cells of Phaffia rhodozyma, results confirmed further that electrostatic interactions were the driving force for the recovery. To elaborate, recoveries up to 97% were obtained with the cationic surfactant CTAB under basic conditions, specifically when the cells of Phaffia rhodozyma were pretreated with NaOH. The use of NaOH had a double effect: it resulted in the higher release of astaxanthin from the cells and it facilitated its subsequent recovery with CTAB. These experiments revealed that the main operating parameters of the separation were the volumetric ratio of CGA to feed or mass ratio of surfactant to total solids in the feed and the initial concentration of total solids in the feed. Specifically, recoveries increased with increased mass ratio and with decreased total solids in the feed. Scale-up with a flotation column The optimum conditions achieved on a small scale (batch), where separation took place in a beaker (the same as for the small scale recovery of proteins, see Fig. 10.3), were applied for the recovery of astaxanthin at larger scale using a flotation column set-up at two different modes (continuous and batch). The set-up for the large scale for the two different modes studied can be seen in Fig. 10.4. Flotation is a separation method that can be used on a large scale but column flotation results in general in lower yields due to the large bubble sizes generated in these columns (Nguyen and Schulze, 2004) and it is more effective for low concentrations of feed and relatively large particle size (mm range) owing to the relative non-turbulent nature of the contact of bubbles and particles (Tasdemir et al., 2007). However, high concentrations of surfactants are necessary and their high cost can restrict their application to the smaller scale (Hashim et al., 1998). CGA, as has already been highlighted above (Table 10.2) have high interfacial area owing to their small size and
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Table 10.4 Summary of the surfactants used and the main outcomes on the recovery of astaxanthin from standard suspensions and suspensions of cells
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cmc mM HLB† Max R % (standard suspensions) Optimum conditions
CTAB (cationic)
SDS (anionic)
Tween 60 (nonionic)
Tween 20 (nonionic)
Span 80 (nonionic)
0.9 – 80%
12 – 55%
0.027 14.9 75%
0.056 16.7 59%
– 4.3 50%
pH = 2 i.s. = 0.055M
Hydrophobic interactions: 20% ethanol 20% Hydrophobic interactions: 20% ethanol
Hydrophobic interactions
Tween 60/Span 80 HLB = 14.024 or HLB = 11.487 – –
Electrostatic interactions: pH = 11 Max R % (raw extract) 97% Optimum conditions Electrostatic interactions: pretreatment with NaOH; 0.2M+homogenisation; clarified Advantages High R% for both natural and synthetic astaxanthin Disadvantages
Toxic Rat: LD50 = 410 mg kg–1
– –
Less toxic than CTAB, Food grade, rat: LD50 = 1200 mg improved biovailability, kg–1 preformulation Low R% Low R% for natural astaxanthin
25% Hydrophobic interactions: Mixture with CTAB nCTAB/nTween2 = 0.1 In H2O + homgenisation Food grade, improved biovailability, preformulation Low R%
Food grade
CGA low stability, low R%
R%, recovery percentage; i.s., ionic strength; LD50, the dose required to kill half the members of a tested population. *CTAB, hexadecyltrimethylammonium bromide; SDS, sodium dodecyl sulfate; Tween 60, polyoxyethylene sorbitan monostearate; Tween 20, polyoxyethylene sorbitan monolaurate; Span 80, sorbitan mono-oleate. † HLB, hydrophile lipophile balance.
Separation of value-added bioproducts by colloidal gas aphrons 301
Surfactant type*
302 Separation, extraction and concentration processes
Flotation column Sample collection CGA feed
CGA generator (a) Feed P2
H4 H3
Flotation column CGA feed
Sample collection
H2 H1
Effluent CGA generator
P1 (b)
P3
Fig. 10.4 Schematic representation of the scale-up of astaxanthin separation with CGA using a flotation column: (a) batch operation, (b) continuous operation.
this makes them more efficient in the adsorption of particles and molecules. Moreover, because of their rheological properties (Table 10.1) they can be easily pumped and they provide increased froth stability which is desirable for flotation processes. Owing to these properties CGA are used for the recovery of a variety of products on a large scale in batch and continuous mode using a flotation column set-up. As shown in Table 10.5, they have been used for the recovery of yeast cells (Hashim et al., 1998; 1995), particles (Mansur et al., 2006; 2004; Waters et al., 2008), oil removal (Watcharasing et al., 2008) and organic dyes from water streams (Pandit and Basu, 2002; Roy et al., 1992); in these studies a range of operating parameters were investigated in order to determine the optimum conditions for the separation. For astaxanthin, the main operating parameter was the volumetric ratio of CGA to feed, which had the same effect in both continuous and batch © Woodhead Publishing Limited, 2010
Table 10.5 Review of applications of CGA in flotation column set-up (c, cationic; a, anionic; n-i, nonionic) (adapted from Dermiki et al., 2010) Surfactant
© Woodhead Publishing Limited, 2010
Separation of CTAB (c) astaxanthin from cells of Phaffia rhodozyma Diesel oil removal from water
Parameters investigated Operation/ dimensions of column Csurf, FCGA, tdrainage, [TS]init, VCGA/Vast, height of the column
Branched alcohol Csurf, tstir, rpm, CNaCl; propoxylate sulfate feed with or without CGA sodium salt (a)
Removal of fine SDS (a) particles
Comparison with conventional flotation, addition of NaCl, pH, particle size
References
Continuous, batch d = 4.5 cm, h = 50 cm
(Dermiki, 2009)
Batch CGA increased froth formation and d = 5 cm, h = 120 stability, enhanced the removal and cm enrichment ratio of oil. Highest oil removal 97%
(Watcharasing et al., 2008)
Batch d = 7 cm, h = 30 cm
SDBS (a), CTAB Surfactant type, Csurf, (c), Tween 20 (n-i) particle size
Batch d = 5 cm h = 100 cm
SDBS (a), CTAB (c)
Batch d = 5 cm, h = 100 cm
Surfactant type, Csurf, FCGA, particle concentration
Main findings
Higher recovery of CuO with CGA (76.5%) than with conventional flotation (58.3%) Addition of NaCl decreases the recovery which implies that electrostatic interactions are important Increase of Csurf up to cmc increases removal efficiency, effect of particle size, no effect of charge of the particles Ionic surfactants: stable dispersions, increased removal efficiency Removal efficiency increases with FCGA and decreases with feed concentration. Better performance of the anionic surfactant due to smaller bubble size of the CGA suspension
(Waters et al., 2008)
(Mansur et al., 2006)
(Mansur et al., 2004)
Separation of value-added bioproducts by colloidal gas aphrons 303
Application
© Woodhead Publishing Limited, 2010
Application
Surfactant
Parameters investigated Operation/ dimensions of column
Main findings
References
Separation of organic dyes
SDBS (a), CTAB (c)
Surfactant type, Csurf, FCGA, CGA diameter and gas hold-up, pH, tres, Csalt
Batch d = 4 cm, h = 60 cm
(Basu and Malpani, 2001)
Surfactant type, FCGA
Batch d @ 8 cm, h @ 100 cm
Removal increases with tres and gas hold-up of CGA. No effects of Csurf above the cmc or Csalt. Higher removal at pH conditions that favour electrostatic interactions Recovery owing to electrostatic interactions, the FCGA affects the removal of dyes
Clarification of palm oil
SDS
Sparging rate of CGA, height of column (tres)
Continuous/ Separation efficiency increases with counter current increasing sparging rate and decreasing d = 5 cm, h = 100 height of the column cm
Treatment of contaminated soil
SDS (a), HTAB (c) Tergitol (n-i) Sapindus mukorossi fruit pericarp
Residence time, type of surfactant, Csurf; CGA and conventional surfactant solution, Csurf, alternating flushing media
d h d h
SDS (a)
Flushing mode, Csurf
d = 5.75 cm h = 10 cm
= = = =
5.75 cm 10 cm 5.75 cm 10 cm
Nonionic more effective. Removal of naphthalene increases with Csurf CGA more effective for removal of hexachlorobenzene (HCB). Recovery increases with Csurf, 24 to 84 times higher than water flushing No effect of alternating media on pressure build-up and removal of HCB CGA more effective. 56% removal of oily waste, 50% removal of transmission fluid CGA more effective in the down flow operation, no effect of Csurf
(Roy et al., 1992) (Hashim et al., 1999)
(Roy et al., 1995) (Kommalapati et al., 1998)
(Roy et al., 1994)
304 Separation, extraction and concentration processes
Table 10.5 Continued
BDHA (c )
Cfeed, FCGA
Continuous/ counter current d = 5 cm h = 58 cm
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pH, Cfeed, FCGA Cfeed, FCGA, operating height
Clarification of suspensions: oil mill effluent, suspension of microalgae, suspension of inorganic minerals
SDS, SDBS (a) CTAC, DTAC, CPC (c)
Surfactant type, CGA height, dispersed phase velocity, cell loading
AOT (a) SDBS (a) BDHA (c) LUX flakes (n-i)
Surfactant type, sparging rate, superficial hydraulic loading rate, air-to-solid ratio, pH, T, flotation cell design
Particle-bubble attachment in yeast flotation follows the Langmuir adsorption mode Separation efficiency = 95%, effect of air flow rate and Cfeed Separation efficiency = 90%, effect of pH at low FCGA
(Hashim et al., 2000) (Hashim et al., 1998) (Hashim et al., 1995)
Effect of FCGA stronger at low Cfeed Concentration gradient along the height of the column, at high Cfeed Continuous/ Better performance of the cationic surfactant (Save and counter current For relatively high dispersed phase Pangarkar d = 3 cm, h = 100 velocity, pH and cell loading have no effect 1995) cm Batch, d = 20 cm V = 5 dm3 Continuous, horizontal flow Continuous/ counter current vertical flow column d = 5 cm, h = 10 cm
Maximum effectiveness of separation at pH (Subramaniam et al., 1990) close to the pK of each surfactant Removal effectiveness increased when CGA were used at the temperature at which they were produced, rather than at high temperatures Effectiveness of solids removal was a function of air-to-solids ratio and of the solids concentration R = 90–94%
Separation of value-added bioproducts by colloidal gas aphrons 305
Yeast cells
306 Separation, extraction and concentration processes experiments. Higher volumetric ratios or mass ratios of surfactant to astaxanthin in the feed resulted in higher amounts of surfactant available to interact with the particles in the feed, leading to increased recovery; the optimum volumetric ratio was 12 where recovery reached 98% and the separation ratio (the astaxanthin concentration ratio in the aphron phase and in the liquid phase) was 10. An increase in the concentration of solids in the feed led to a lower separation efficiency (Dermiki et al., 2010), which is in agreement with other studies on recovery of yeast cells (Hashim et al., 1998) and particles. This is because of saturation of the aphron phase with particles. For the batch experiments, flow rate of CGA had a significant effect on the selectivity of the process. As the flow rate increased from 70 to 150 ml min–1 selectivity reached a maximum and further increase resulted in reduced selectivity. This is because the flow rate caused changes in contact time and CGA stability which had both, a positive and a negative effect on selectivity. Increased flow rate led to increased turbulence and decreased drainage. Stability must be high enough to ensure attachment of astaxanthin but not too high so that the liquid within the bubbles is drained and cell particles partition to the bulk liquid phase (Dermiki et al., 2010). In continuous operation, flow rate and height or the residence time of CGA in the column were the main operating parameters. Hence, separation increased with increasing flow rate as for the batch operation mode. For a specific flow rate, separation increased at higher levels of the column as this corresponded to longer residence times of CGA. However, the optimum flow rate for separation did not lead to the highest selectivity, which was also in agreement with the results in the batch experiments. The experiments with astaxanthin showed that because the astaxanthin particles are significantly smaller than the cells, a different mechanism is responsible for their separation in a flotation column: adsorption of small particles to the bubbles is mainly driven by electrostatic interactions, whereas attachment of the cell/cell aggregates to the bubbles occurs mainly because of non-specific interfacial forces such as buoyancy forces (Klimpel, 1998). The studies described showed that CGA produced from low surfactant concentrations can be used for the recovery of bioproducts and for astaxanthin recoveries up to 98% could be achieved using CGA generated from CTAB solutions of 0.8 mM, which is below the cmc of CTAB (Dermiki et al., 2010) Optimum recoveries of astaxanthin were obtained with the cationic surfactant CTAB, which at high concentrations could be toxic. However, its removal was possible by ultrafiltration with membranes of regenerated cellulose with 50 kDa molecular weight cut off, at low pH where electrostatic interactions between astaxanthin and CTAB were not promoted (Dermiki, 2009). Furthermore, by using diafiltration under the same pH conditions as before, it was possible to obtain CTAB concentrations in the retentate that were below the toxic limit. Thus, the final product was not considered toxic and the recovered surfactant in the permeate could be recycled and reused.
© Woodhead Publishing Limited, 2010
Separation of value-added bioproducts by colloidal gas aphrons 307 For astaxanthin the use of nonionic surfactants was investigated because they offer some important advantages. The surfactants that were investigated here are food-grade surfactants therefore their removal is not necessary at the final stage of the process. Moreover, these surfactants enhance the bioavailability of hydrophobic compounds such as astaxanthin. It must be stressed that astaxanthin is already available in emulsion form; therefore the use of nonionic surfactants can integrate the recovery and formulation steps. As shown in Table 10.4, high recoveries were obtained for synthetic astaxanthin particles in the standard suspension experiments. However, these high recoveries were not repeated for astaxanthin from the cells of Phaffia rhodozyma. These differences in recovery can be explained on the basis of the differences in size between the synthetic astaxanthin (>1 mm) and natural astaxanthin (<0.2 mm) and on the differences in stability between CGA generated from nonionic and ionic surfactants (Dermiki, 2009). Nonionic surfactants form less stable CGA dispersions that are not appropriate for the recovery of small particles such as the natural astaxanthin particles. Despite the fact that nonionic surfactants used on their own showed relatively low recoveries of natural astaxanthin particles, recovery and selectivity increased when the nonionic surfactants were combined with the cationic CTAB (Dermiki, 2009). The use of nonionic surfactants can be advantageous and should be investigated further for the recovery of astaxanthin from other sources, such as the microalgae Haematococcus pluviallis. In this instance, astaxanthin is esterified therefore hydrophobic interactions with the nonionic surfactant are important and promoting these could lead to high recoveries.
10.4 Feasibility of industrial application of CGA The case studies presented in this chapter show that CGA can be applied to the separation of a range of value-added products from waste streams, as well as from fermentation broths. In this section the feasibility of CGA for these industrial applications is explored. An economical evaluation of CGA separation and other conventional separations used in the recovery of astaxanthin at industrial scale was carried out. The economical evaluation of CGA was compared with that of solvent extraction (SE) and supercritical carbon dioxide extraction (SCDE) in terms of operating cost and cost of equipment (Dermiki, 2009). Process economics were evaluated using the software SuperPro Designer (Intelligen, US). One of the main outcomes of this study was that when comparing the purchase cost of equipment, CGA was the most expensive. This can be attributed to the high cost of diafiltration for the removal of surfactant; if this step was removed, the equipment cost was about the same as that of supercritical carbon dioxide extraction and lower than that of solvent extraction. However, another interesting outcome was that the latter had the © Woodhead Publishing Limited, 2010
308 Separation, extraction and concentration processes highest total operational cost, mainly owing to the high cost of raw materials, followed by supercritical carbon dioxide extraction. Although surfactants are also very expensive, low concentrations are required to generate CGA and surfactants cost less than solvents. Furthermore, the cost of utilities, which includes the cost of electricity, is high for CGA extraction but it is not higher than that of SCDE despite the high energy input requirement for CGA generation. In SCDE, the electricity cost for the compression of carbon dioxide increases the overall cost of utilities. It was also interesting to find that the cost of waste disposal was the highest for SE, very similar to that of SCDE and the lowest was for CGA. In summary, although this was an approximate estimation of the economics of a CGA separation and, therefore, several assumptions had to be made, it is still a very valuable evaluation as it highlights the strengths and weaknesses of CGA separations in relation to other conventional industrial separations. From this, it can be concluded that if CGA are applied following an integrated approach, i.e. if removal of surfactant is not required, this separation could be more advantageous than others as it is more cost effective. Furthermore, CGA separation is a more environmentally friendly separation process than solvent extraction and also uses less hazardous equipment than supercritical carbon dioxide extraction.
10.5 Future trends As shown above one of the drawbacks of the use of CGA to separate bioproducts is the cost of the surfactant and its toxicity. Removal of the surfactant can be achieved in some instances as shown in the astaxanthin study by membrane filtration/diafiltration. Although this allows recycling of the surfactant, it leads to an overall increase in the operating cost thus its application becomes less attractive. For some applications nonionic surfactants can be applied in place of the more toxic ionic surfactants and thus further process integration can be achieved as recovery of the product in an aqueous solution of a nonionic surfactant may result also in its preformulation. Therefore, the use of nonionic surfactants should be explored further. We investigated the use of nonionic surfactants for the recovery of astaxanthin but low recoveries were achieved (<20%). However, recovery and selectivity increased slightly when the nonionic surfactant was combined with the cationic surfactant CTAB (Dermiki, 2009). This is a way of decreasing the amount of the cationic surfactant. An interesting alternative to the toxic surfactants are the biosurfactants because these are nontoxic, biodegradable and most importantly they can be produced from several inexpensive waste sources making them an alternative for many synthetic surfactants (Singh et al., 2007). For CGA to be competitive against other separations, the integrative approach has to be possible as shown for astaxanthin. The application of CGA directly to the suspension of cells resulted in high recoveries and partial selective separation. © Woodhead Publishing Limited, 2010
Separation of value-added bioproducts by colloidal gas aphrons 309 As demonstrated also by the economic evaluation above, reduction in the number of process steps will reduce the overall cost. Similarly, Zidehsarei et al. (2009) generated CGA together with the feed suspension, which consisted of a solid fermentation suspension with 50% wet bran, Aspergillus niger and the product, glucoamylase. This method enabled them to simultaneously extract (from the solid substrate and biomass) and recover the enzyme in the drained liquid with recovery of 91% and specific activity of 10. In addition, they found that using aluminium sulfate as a flocculating agent they achieved improved separation of the solids in the CGA and improved drainage, which resulted in higher purity. So these applications show that CGA could be advantageous as a bioseparation for some of the value-added products and may be able to find industrial applications where high yield rather than purity is a requirement. Adsorption-based processes such as chromatography are more selective and therefore CGA cannot compete with these. Other interesting applications of CGA include environmental applications such as the remediation of soil contaminated with diesel oil (Couto et al., 2009), removal of organic dyes from wastewaters (Roy et al., 1992) and the removal of pyrene from soil with biodegradable surfactants (BioSolve and BioNex) (Boonamnuayvitaya et al., 2009). The latter shows that CGA can be generated with biodegradable surfactants and thus they are suitable for environmental applications such as remediation. This is also shown in the work by Park and co-workers (2009) who applied CGA made with a biodegradable surfactant saponin, to biodegradation of phenanthrene. They used CGA to enhance oxygen transfer during the fermentation process. Furthermore, other applications of CGA include their use to enhance ultrasound imaging owing to their reflection properties (Wheatley et al., 2006), the application as drug delivery agents using the biocompatible synthetic polymer poly(vinyl alcohol) (PVA) (Cavalieri et al., 2005) and their application as a template for the formation of tungsten oxide nanorods (Abdullah et al., 2006).
10.6 Sources of further information and advice Rosen M J (2004) Surfactants and interfacial phenomena, New Jersey, John Wiley and Sons., Inc. Sebba F (1987) Foams and biliquid foams–aphrons. Chichester, John Wiley & Sons Ltd. Nguyen A V and Schulze H J (2004) Colloidal science of flotation, New York, Marcel Dekker, Inc. United States Code of Federal Regulations (CFR) (2009) 21CFR172–Part 172 ‘Food additives permitted for direct addition to food’ Available from: http://frwebgate.access.gpo.gov/cgi-bin/multidb.cgi.
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310 Separation, extraction and concentration processes
10.7 References Abdullah SF, Radiman S, Abd. Hamid MA and Ibrahim NB (2006), ‘Effect of calcination temperature on the surface morphology and crystallinity of tungsten (VI) oxide nanorods prepared using colloidal gas aphrons method’, Colloids and Surfaces A: Physicochemical and Engineering Aspects, 280(1–3), 88–94. Alves RW, de Souza AAU, de Souza SMD and Jauregi P (2006) ‘Recovery of norbixin from a raw extraction solution of annatto pigments using colloidal gas aphrons (CGAs)’, Separation and Purification Technology, 48(2), 208–213. Amiri MC and Valsaraj KT (2004), ‘Effect of gas transfer on separation of whey protein with aphron flotation’, Separation and Purification Technology, 35, 161–167. Amiri MC and Woodburn ET (1990), ‘A method for the characterization of colloidal gas aphrons’, Transactions IChemE, 68, 154–160. Basu S and Malpani PR (2001), ‘Removal of methyl orange and methylene blue from water using colloidal gas aphrons – effect of process parameters’ Separation Science and Technology, 36, 2997–3013. Bhatia D, Goel G, Bhimania SK and Bhaskarwar AN (2005), ‘Characterization and drainage kinetics of colloidal gas aphrons’, AIChE Journal, 51, 3048–3058. Boonamnuayvitaya V, Jutaporn P, Sae-ung S and Jarudilokkul S (2009), ‘Removal of pyrene by colloidal gas aphrons of a biodegradable surfactant’, Separation and Purification Technology, 68, 411–416. Bredwell MD, Telcenhoff MD and Worden RM (1995), ‘Formation and coalescence properties of microbubbles’, Applied Biochemistry and Biotechnology, 51/52, 501–508. Bredwell M and Worden R (1998), ‘Mass-transfer properties of microbubbles. 1. Experimental studies’, Biotechnology Progress, 14, 31–38. Brookey TF and Tommy F (2004) ‘Aphron-containing well drilling and servicing fluids’. US patent 6716797. Cavalieri F, El Hamassi A, Chiessi E and Paradossi G (2005), ‘Stable polymeric microballoons as multifunctional device for biomedical uses: synthesis and characterization’, Langmuir, 21, 8758–8764. Chaphalkar PG, Valsaraj KT and Roy D (1993), ‘A study of the size distribution and stability of colloidal gas aphrons using a particle size analyser’, Separation Science and Technology, 28, 1287–1302. Chu W (2003), ‘Remediation of contaminated soils by surfactant-aided soil washing’, Practice Periodical of Hazardous, Toxic, and Radioactive Waste Management, 7, 19–24. Couto HJB, Massarani G, Biscaia Jr EC and Sant’Anna Jr GL (2009), ‘Remediation of sandy soils using surfactant solutions and foams’, Journal of Hazardous Materials, 164, 1325–1334. Dai Y and Deng T (2003), ‘Stabilization and characterization of colloidal gas aphron dispersions’, Journal of Colloid and Interface Science, 261(2), 360–365. Dai Y, Deng T, Wang J and Xu K (2004), ‘Enhancement of oxygen gas–liquid mass transfer with colloidal gas aphron dispersions’, Colloids and Surfaces A: Physicochemical and Engineering Aspects, 240, 165–171. Dermiki M, Gordon MH and Jauregi P (2008), ‘The use of colloidal gas aphrons as novel downstream processing for the recovery of astaxanthin from cells of Phaffia rhodozyma’, Journal of Chemical Technology and Biotechnology, 83, 174–182. Dermiki M (2009), ‘Recovery of astaxanthin using colloidal gas aphrons’, PhD thesis, Reading University, UK. Dermiki M, Bourquin A-L and Jauregi P (2010), ‘Separation of astaxanthin from cells of Phaffia rhodozyma using colloidal gas aphrons (CGA) in a flotation column’, Biotechnology Progress, 26(2), 477–487. Dermiki M, Gordon MH and Jauregi P (2009), ‘Recovery of astaxanthin using colloidal
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Separation of value-added bioproducts by colloidal gas aphrons 311 gas aphrons (CGA): a mechanistic study’, Separation and Purification Technology, 65, 54–64. Feng W, Singhal N and Swift S (2009), ‘Drainage mechanism of microbubble dispersion and factors influencing its stability’, Journal of Colloid and Interface Science, 337, 548–554. Fernandes S, Hatti-Kaul R and Mattiasson B (2002a), ‘Selective recovery of lactate dehydrogenase using affinity foam’, Biotechnology and Bioengineering, 79, 472– 480. Fernandes S, Mattiasson B and Hatti-Kaul R (2002b), ‘Recovery of recombinant cutinase using detergent foam’, Biotechnology Progress, 18, 116–123. Fuda E, Bhatia D, Pyle DL and Jauregi P (2005), ‘Selective separation of b-lactoglobulin from sweet whey using CGAs generated from a cationic surfactant CTAB’, Biotechnology and Bioengineering, 90, 532–542. Fuda E, Jauregi P and Pyle DL (2004), ‘Recovery of lactoferrin and lactoperoxidase from sweet whey using colloidal gas aphrons (CGAs) generated from an anionic surfactant, AOT’, Biotechnology Progress, 20, 514–525. Fuda E and Jauregi P (2006) ‘An insight into the mechanism of protein separation by colloidal gas aphrons (CGA) generated from ionic surfactants’, Journal of Chromatography B, 843, 317–326. Hashim MA, Dey A, Hasan S and Sen Gupta B (1999), ‘Mass transfer correlation in flotation of palm oil by colloidal gas aphrons’, Bioprocess Engineering, 21, 401–404. Hashim MA, Kumar SV and SenGupta B (2000), ‘Particle-bubble attachment in yeast flotation by colloidal gas aphrons’, Bioprocess and Biosystems Engineering, 22, 333–336. Hashim MA, Sen Gupta B, Kumar VS, Lim R, Lim SE and Lim C (1998), ‘Effect of air to solid ratio in the clarification of yeast by colloidal gas aphrons’, Journal of Chemical Technology & Biotechnology, 71, 335–339. Hashim MA, Sen Gupta B and Subramaniam MB (1995), ‘Investigations on the flotation of yeast cells by colloidal gas aphrons (CGA)’, Bioseparation, 5, 167–173. Hussein G, Sankawa U, Goto H, Matsumoto K and Watanabe H (2006), ‘Astaxanthin, a carotenoid with potential in human health and nutrition’, Journal of Natural Products, 69, 443–449. Jarudilokkul S, Rungphetcharat K and Boonamnuayvitaya V (2004), ‘Protein separation by colloidal gas aphrons using nonionic surfactant’, Separation and Purification Technology, 35, 23–29. Jauregi, P (1997), ‘Colloidal gas aphrons (CGA) a novel approach to protein recovery’, PhD thesis, Reading University, UK. Jauregi P, Gilmour S and Varley J (1997), ‘Characterisation of colloidal gas aphrons for subsequent use for protein recovery’, Chemical Engineering Journal, 65, 1–11. Jauregi P, Mitchell GR and Varley J (2000), ‘Colloidal gas aphrons (CGA): dispersion and structural features’, AIChE Journal, 46, 24–36. Jauregi P and Varley J (1999), ‘Colloidal gas aphrons: potential applications in biotechnology’, Trends in Biotechnology, 17, 389–395. Kaster JA, Michelsen DL and Velander WH (1990), ‘Increased oxygen transfer in a yeast fermentation using a microbubble dispersion’, Applied Biochemistry and Biotechnology, 24/25, 469–484. Klimpel R (1998), ‘Introduction to solid–solid separation of fine particles’, Florida, NSF Engineering Research Center for Particle Science and Technology. Kommalapati RR, Valsaraj KT, Constant WD and Roy D (1998), ‘Soil flushing using colloidal gas aphron suspensions generated from a plant-based surfactant’, Journal of Hazardous Materials, 60, 73–87. Larmignat S, Vanderpool D, Lai HK and Pilon L (2008), ‘Rheology of colloidal gas aphrons (microfoams)’, Colloids and Surfaces A: Physicochemical and Engineering Aspects, 322, 199–210.
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312 Separation, extraction and concentration processes Mansur EHA, Wang Y and Dai Y (2004), ‘Separation of fine particles by using colloidal gas aphrons’, Chinese Journal of Chemical Engineering, 12, 286–289. Mansur EHA, Wang Y and Dai Y (2006), ‘Removal of suspensions of fine particles from water by colloidal gas aphrons (CGAs)’, Separation and Purification Technology, 48, 71–77. Matsushita K, Mollah AH, Stuckey DC, del Cerro C and Bailey AI (1992), ‘Predispersed solvent extraction of dilute products using colloidal gas aphrons and colloidal liquid aphrons: aphron separation, stability and size’, Colloids and Surfaces, 69, 65–72. Monsalve-Gonzalez A and Ochomogo M (2009) ‘Natural flavour enhancement compositions for food emulsions’, US patent 0196972. Moshkelani M and Amiri MC (2008), ‘Electrical conductivity as a novel technique for characterization of colloidal gas aphrons (CGA)’, Colloids and Surfaces A: Physicochemical and Engineering Aspects, 317, 262–269. Nguyen AV and Schulze HJ (2004), ‘Industrial Applications’, in Nguyen AV and Schulze HJ, Colloidal Science of Flotation, New York, Marcel Dekker, 813–842. Noble M, Brown A, Jauregi P, Kaul A and Varley J (1998), ‘Protein recovery using gas–liquid dispersions’, Journal of Chromatography B, 711, 31–43. Noble MJ and Varley J (1999), ‘Colloidal gas aphrons generated from the anionic surfactant AOT for the separation of proteins from aqueous solution’, Journal of Chemical Technology and Biotechnology, 74(3), 231–237. Noel J, Ales P and Tanner R (2002), ‘Foam fractionation of a dilute solution of bovine lactoferrin’, Applied Biochemistry and Biotechnology, 98–100, 395–402. O’Connell PJ and Varley J (2001), ‘Immobilization of Candida rugosa lipase on colloidal gas aphrons (CGAs)’, Biotechnology and Bioengineering, 74, 264–269. Pandit P and Basu S (2002), ‘Removal of organic dyes from water by liquid–liquid extraction using reverse micelles’, Journal of Colloid and Interface Science, 245, 208–214. Park JY, Choi YJ, Moon S, Shin DY and Nam K (2009), ‘Microbubble suspension as a carrier of oxygen and acclimated bacteria for phenanthrene biodegradation’, Journal of Hazardous Materials, 163, 761–767. Rosen M (2004), ‘Foaming and antifoaming by aqueous solutions of surfactants’, Surfactants and interfacial phenomena, New Jersey, John Wiley & Son. Roy D, Kommalapati RR, Valsaraj KT and Constant WD (1995), ‘Soil flushing of residual transmission fluid: application of colloidal gas aphron suspensions and conventional surfactant solutions’, Water Research, 29, 589–595. Roy D, Valsaraj KT, Constant WD and Darji M (1994), ‘Removal of hazardous oily waste from a soil matrix using surfactants and colloidal gas aphrons under different flow conditions’, Journal of Hazardous Materials, 38, 127–144. Roy D, Valsaraj KT and Kottai SA (1992), ‘Separation of organic dyes from wastewater by using colloidal gas aphrons’, Separation Science and Technology, 27, 573–588. Save SV and Pangarkar VG (1994), ‘Characterisation of colloidal gas aphrons’, Chemical Engineering Communications, 127, 35–54. Save SV and Pangarkar VG (1995), ‘Harvesting of Saccharomyces cerevisiae using colloidal gas aphrons’, Journal of Chemical Technology and Biotechnology, 62, 192–199. Sebba F (1972), ‘Biliquid foams – a preliminary report’ Journal of Colloid and Interface Science, 40, 468–474. Sebba F (1985), ‘An improved generator for micron-sized bubbles’, Chemistry and Industry, 3, 91–92. Sebba F (1987), Foams and Biliquid Foams–Aphrons. Chichester, John Wiley & Sons Ltd. Singh A, Van Hamme JD and Ward OP (2007), ‘Surfactants in microbiology and biotechnology: Part 2. Application aspects’. Biotechnology Advances, 25, 99–121. Spigno G, Dermiki M, Pastori C, Casanova F and Jauregi P (2010), ‘Recovery of gallic acid with colloidal gas aphrons generated from a cationic surfactant’, Separation and Purification Technology, 71, 56–62. © Woodhead Publishing Limited, 2010
Separation of value-added bioproducts by colloidal gas aphrons 313 Spigno G and Jauregi P (2005), ‘Recovery of gallic acid with colloidal gas aphrons (CGA)’, International Journal of Food Engineering, 1, 1–10. Subramaniam MB, Blakebrough N and Hashim MA (1990), ‘Clarification of suspensions by colloidal gas aphrons’, Journal of Chemical Technology and Biotechnology, 48, 41–60. Tasdemir A, Tasdemir T and Öteyaka B (2007), ‘The effect of particle size and some operating parameters in the separation tank and the downcomer on the Jameson cell recovery’, Minerals Engineering, 20, 1331–1336. Tseng H, Pilon L and Warrier GR (2006), ‘Rheology and convective heat transfer of colloidal gas aphrons in horizontal mini-channels’, International Journal of Heat and Fluid Flow, 27, 298–310. Watcharasing S, Angkathunyakul P and Chavadej S (2008), ‘Diesel oil removal from water by froth flotation under low interfacial tension and colloidal gas aphron conditions’, Separation and Purification Technology, 62, 118–127. Waters KE, Hadler K and Cilliers JJ (2008), ‘The flotation of fine particles using charged microbubbles’, Minerals Engineering, 21, 918–923. Wheatley M, Forsberg F, Dube N, Patel M and Oeffinger B (2006), ‘Surfactant-stabilized contrast agent on the nanoscale for diagnostic ultrasound imaging’, Ultrasound in Medicine and Biology, 32, 83–93. Worden RM and Scranton AB (2000), ‘Method for forming reversible colloidal gas or liquid aphrons and compositions produced’, US patent 6022727. Xu Q, Nakajima M, Ichikawa S, Nakamura N and Shiina T (2008), ‘A comparative study of microbubble generation by mechanical agitation and sonication’, Innovative Food Science & Emerging Technologies, 9, 489–494. Yan Y-L, Qu C-T, Zhang N-S, Yang Z-G and Liu L (2005), ‘A study on the kinetics of liquid drainage from colloidal gas aphrons (CGAs)’, Colloids and Surfaces A: Physicochemical and Engineering Aspects, 259, 167–172. Zhao J, Pillai S and Pilon L (2009), ‘Rheology of colloidal gas aphrons (microfoams) made from different surfactants’, Colloids and Surfaces A: Physicochemical and Engineering Aspects, 348, 93–99. Zidehsaraei AZ, Moshkelani M and Amiri MC (2009), ‘An innovative simultaneous glucoamylase extraction and recovery using colloidal gas aphrons’, Separation and Purification Technology, 67, 8–13.
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11 Membrane bioreactors and the production of food ingredients M.-P. Belleville, D. Paolucci-Jeanjean and G. M. Rios, European Institute of Membranes, France
Abstract: Membrane bioreactors (MBRs) and their potential uses in food processing and food ingredients production are discussed, where ‘membrane bioreactor’ applies to any system constituted from a reactor working with enzymes or whole cells as the catalyst and from a membrane for separation and/or contacting purposes. The various types of MBR are presented and their advantages as well as their drawbacks are discussed. Several applications of MBRs in different food areas are then presented. Key words: membrane bioreactors, enzymatic membrane reactors, whole cell membrane reactors, bioproduction, food processing.
11.1 Introduction With increasing awareness of environmental and cost issues, biotransformations are gaining ground rapidly owing to the advantages that they offer over conventional technologies. Enzyme catalyzed reactions occur with high rates at room temperature (thermal degradation of labile compounds is avoided) with a minimum use of chemicals and a reduced number of reaction steps by avoiding by-products. Furthermore, owing to recent food scares, consumers’ fears of artificial products have increased the interest in biotransformation processes, which lead to products regarded as natural ingredients. Generally, bioconversions are carried out in classical batch reactors, which are easy to implement at any scale but present several disadvantages, in particular on the industrial scale: low efficiency owing to start up and shut down procedures, high labour costs and great variability of product quality owing to batch-to-batch variations. Moreover, at the end of the reaction,
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Membrane bioreactors and the production of food ingredients 315 biocatalysts are removed even if they are still active entailing high processing costs in particular when enzymes are used as catalyst. To reduce enzyme cost, enzyme immobilization on solid supports has been widely used. Thus, at the end of reaction, the catalyst can be easily separated from the reaction medium, cleaned and reused in further reactions. Moreover immobilized enzymes can be used in semicontinuous or continuous processes at the industrial stage because immobilization improves their stability (Iyer and Ananthanarayan, 2008) as well as other properties such as specificity and selectivity (Mateo et al., 2007). The operational costs are then reduced. The membrane reactor that associates reaction and membrane separation can represent an attractive alternative approach to classical methods of biocatalyst immobilization. In such a reactor, the membrane ensures the complete rejection of the enzyme or cells in order to maintain the full activity inside the reacting volume without the need for other immobilization techniques while the products are continuously extracted from the medium, thus reducing their inhibitory effects on the reaction rate. The first studies on membrane reactors applied to biological or food processes, here referred to as membrane bioreactors (MBRs), were carried out in the 1970s and since then numerous reviews have been published (Belfort, 1989; Cheryan and Mehaia, 1986; Giorno and Drioli, 2000; Prazeres and Cabral, 1994; Rios et al., 2004; Sanchez Marcano and Tsotsis, 2002). This chapter focuses on membrane bioreactors and their potential uses in food processing and food ingredients production. Firstly, the various types of MBR are presented and their advantages as well as their drawbacks are discussed. Then, several applications of MBRs in different food areas are given.
11.2 Membrane bioreactors for the production of food ingredients Membrane bioreactors correspond to the association of a reactor involving enzymes or whole cells as catalyst with a membrane separation unit. However, MBRs can be classified into two types depending on the role played by the membrane. The first type, also called free enzymes or cells membrane bioreactors (FEMBR or FCMBR) concerns MBRs involving free enzymes or cells and in that instance the function of the membrane is to retain the biocatalysts inside the reactor throughout the process. In the second type, the biocatalyst is immobilized on the membrane surface or within its pore structure and the reaction occurs at the outside or internal surface of the material during the membrane transfer.
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316 Separation, extraction and concentration processes 11.2.1 Free enzymes or free cells membrane bioreactors (FEMBRs or FCMBRs) The first configuration (Fig. 11.1) still remains the most popular and simplest one. In the continuous mode, a stirred tank reactor continuously fed with fresh substrate solution is associated with a membrane separation unit. In order to limit polarization phenomena and membrane fouling, the reaction medium generally flows tangentially along the membrane before being recycled into the reactor. Such reactors are also referred to as continuous recycle membrane bioreactors (CRMBR). The choice of the membrane is essential as regards performance; the membrane molecular weight cut off (MWCO) should be chosen in order to ensure the retention of enzymes or cells as well as the substrate; but the membrane pores must be large enough to enable the product to pass through the membrane module. As many enzymes have a molecular weight of 10 to 80 kDa, ultrafiltration membranes with a molecular cut-off between 1 and 100 kDa are generally used in FEMBR. For FCMBRs, microfiltration membranes are generally preferred. Chèze-Lange et al. (2002) reported that pore diameters should be below 1.4 mm, in order to prevent cell leakage. For the membrane material, a wide range of polymeric membranes, including polysulfone, polyethersulfone, polyamide and polypropylene, is used as well as inorganic membranes. Although mineral membranes are more expensive than organic ones, they are often more attractive because they can endure higher temperatures and pressures as well as strong chemical treatments for regeneration or even sterilization. Particular attention should be paid to the selection of the membrane material because electrostatic or hydrophobic interactions between the enzymes or cells and the membrane could affect the process performance. In particular, enzyme or cell adsorption leads to a reduction of permeation flux owing to membrane fouling. Moreover, as reported by Bódalo et al. (2004) and Paolucci-Jeanjean et al. (2001), enzyme
Substrate
Products
Membrane unit
Biocatalysts
Reactor unit
Fig. 11.1 Continuous recycle membrane bioreactor (CRMBR).
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Membrane bioreactors and the production of food ingredients 317 adsorption onto the membrane surface may induce a catalytic efficiency decrease in FEMBRs. In order to improve the control of fouling Kwon et al. (2006) suggest the use of submerged membrane reactor (Fig. 11.2). This MBR configuration, more generally called immersed membrane bioreactor (iMBR), is widely used for aerobic wastewater treatment (Judd, 2008). In practice, the membrane is placed inside the bioreactor and permeate is obtained thanks to suction force or to the transmembrane pressure owing to hydrostatic pressure. In such reactors, fouling is avoided by the use of coarse-bubble aeration which generates bubbles close to the membrane surface. iMBRs are also more efficient in regard to cells recycling and power consumption. Although 30 to 50% of the energy demand is caused by aeration of the membrane, high air fluxes promote higher permeate flux and thus lower the required membrane areas. Furthermore, the use of intermittent air flux or optimized design of the membrane module can reduce the energy consumption (Judd, 2008). In a third type of FEMBRs or FCMBRs: hollow-fibre MBRs, biocatalysts (enzymes or cells) are retained on the shell side whereas substrate flows along the lumen side of the membrane (Krastanov et al., 2007; Novalin et al., 2005). In such reactors, mass transfer resistances are high; the substrate has to be transferred across the membrane to the shell side where the biocatalyst is located and then the product has to be transferred back to the lumen side of the membrane. In addition to the reactors described above, it is also possible to couple a bioreactor to a pervaporation unit. For enzymatic ester synthesis, this coupling enables an effective in situ removal of water (a by-product of the reaction) through hydrophilic pervaporation membranes. In this way, the reaction equilibrium is displaced in favour of synthesis and the conversion rate is increased (Bartling et al., 2001; Won et al., 2006; Ziobrowski et al., 2009). In a recent review, Vane (2005) detailed the interest of such coupling for product recovery from biomass fermentation processes. The pervaporation unit can be directly associated with the reactor except for reactions involving thermosensitive cells or enzymes. In practice, the performance is improved Products Substrate
Membrane Cells Reactor
Fig. 11.2 Submerged or immersed membrane bioreactor (iMBR).
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318 Separation, extraction and concentration processes when the system operates at high temperature. The non-thermoresistant biocatalysts need to be removed before being subjected to denaturating conditions. In that instance, the pervaporation module can be fed with the permeate obtained from a CRMBR (Del Amor Villa and Wichmann, 2005). To avoid product or by-product inhibition, pervaporation can also be replaced by perstraction or electrodialysis. Qureshi and Maddox (2005) combined a batch reactor with a perstraction unit in order to produce butanol from concentrated lactose-whey. Valadez-Blanco et al. (2008) associated a bioreactor with a nanofiltration membrane contactor in order to extract continuously the R-citronellol produced by baker’s yeast from geraniol in hexane. This MBR configuration avoided the formation of emulsions, thus reducing downstream separation and enabling increased substrate loadings. Meynial-Salle et al. (2008) developed an integrated membrane–bioreactor–electrodialysis system that permits the bioproduction of succinic acid at high concentration, productivity and yield using Anaerobiospirillum succiniciproducens. Compared with classical reactor configurations such as batch, fixed or fluidized-bed reactors, free enzyme or cell MBRs offer some advantages. They permit a continuous operating mode and the high concentration of biocatalyst as well as the continuous removal of inhibitors. This entails high production rates, which ensure the economic viability of the process. In addition, as reaction and separation zones are placed in series, they can be dealt quite independently in order to optimize the performance of the whole process. Consequently, production can be optimized by acting separately on temperature, pH, substrate and biocatalyst concentrations, fluid velocity, individual control of hydraulic and biomass residence times, pressure, reactor volume or membrane surface. Unfortunately, such MBRs also present some drawbacks. The first problem relates to the change in biocatalyst activity. The loss of activity is essentially owing to adsorption onto membrane surface as well as mechanical stress, which entails enzyme and cell deactivation. However, for reactors involving cells, a part of this loss may be caused by nutrient limitation or toxin accumulation in the broth. The second main drawback is membrane fouling, which reduces the permeate flux and thus the cost-efficiency of the process. In a recent review, Meng et al. (2009) present the recent and current developments concerning the fouling behaviour, fouling factors and fouling control strategies in MBRs. Finally, if the use of FEMBR is particularly recommended when enzyme and substrate are larger than products (as in hydrolysis reactions), they are not adapted for carrying out synthesis when substrates and products show similar sizes. In this latter instance, it is better to use a reactor where the biocatalysts are immobilized on or within the membrane. Such membranes, also named ‘active membranes’, are at the heart of the so-called immobilized enzyme membrane bioreactor (IEMBR) described here below.
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Membrane bioreactors and the production of food ingredients 319 11.2.2 Immobilized enzyme membrane bioreactor (IEMBR) There are two modes for operating active membranes and associated IEMBRs as presented in Fig. 11.3. In the first mode (Fig. 11.3a), the substrate solution flows through the membrane to the biocatalyst as a result of a transmembrane pressure. The reaction occurs at the wall and the product is recovered in the permeate. The membrane is a quite specific macrosystem resulting from the assembly of swarms of microsystems (the pores), which can be regarded as microreactors. In such reactors, the mass transfer path is reduced, the contact between substrate molecules and catalysts are thus favoured. A precise control of the reaction with minimized substrate and catalyst losses, faster reactions and higher yields and cleaner products can be expected (Rios et al., 2004). In addition as the membrane thickness is much smaller than the bed length, the pressure drop and energy costs are dramatically reduced compared with fixed or fluidized bed reactors. In the second mode (Fig. 11.3b), reaction takes place at a phase-contacting interface within the membrane material. Two solutions (the substrate solution and an extracting solution that is not miscible with the substrate one) flow along each membrane side. The substrate diffuses from the feed solution, reacts with the biocatalyst and the product diffuses towards the extracting solution with which it presents a very high affinity. This biphasic reactor is particularly attractive when products and substrate show different solubilities in water and organic media as is the case in lipid hydrolysis (Knezevic et al., 2004; Merçon et al., 2000; Wang et al., 2008) or lipase-catalyzed kinetic resolution of racemic solution (Long et al., 2005; Ong et al., 2008; Wang et al., 2007). The reaction can occur without requiring emulsion formation and the product can be obtained in a single phase. However, compared with FEMBRs, the performance of such a reactor is limited by a high mass transfer resistance; indeed mass transport occurs by diffusion. Whatever the IEMBR considered, one of the important considerations is the proper incorporation of the active catalyst on or within the membrane. The three main types of active membrane preparation are shown in Fig. 11.4.
m zy
Substrate
En
Permeate
Aqueous phase
Enzyme
Feed
e
Retentate
Organic phase
Products
(a)
(b)
Fig. 11.3 Immobilized enzymatic membrane bioreactors (IEMBRs): monophasic reactor (a); biphasic reactor (b).
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320 Separation, extraction and concentration processes The biocatalyst can be (1) entrapped within the membrane structure, (2) immobilized by gelification on the membrane surface or (3) attached through covalent or noncovalent binding at the membrane surface. Entrapment within the polymeric structure can be achieved by mixing the enzyme solution with the polymeric solution used for membrane formation. The biocatalyst can be simply physically entrapped or it can be covalently linked to the polymer matrix to avoid enzyme leakage (Kanwar and Goswami, 2002; Tan et al., 2002). It is also possible to add particulate support (i.e. active carbon) previously loaded with enzymes in the polymeric solution before casting the polymeric film (Torras et al., 2008). However, the most widely used method to achieve enzyme entrapment is filtration. The enzyme solution is filtered from the support to the separating layer in order to retain the biocatalyst in the porosity of the membrane support (Sousa et al., 2001; Wang et al., 2008; Xu et al., 2006). This is a very simple procedure that leads to high enzyme loading but also has high leaching risks. Such active membranes are generally used when the feed solution permeates from the outside to the inside of the membrane or for the biphasic reactor. In this latter instance, the low solubility of proteins in organic solvent prevents their desorption. In order to limit the risk of enzyme leakage and improve the enzyme stability, Hilal et al. (2004) suggest embedding crosslinked enzyme aggregates (CLEAs) within membrane porosity. The CLEAs are formed inside the membrane pores previously filled with enzyme solution by precipitation using organic solvent with simultaneous crosslinking by glutaraldehyde. It is worth noting that active membranes elaborated by entrapment are more convenient for the reaction-limited regime rather than for a diffusion-limited regime. Active membranes can also be obtained by enzyme gelification on the membrane surface. The enzyme solution is filtered from the active side to the porous support on an ultra- or microfiltration membrane and the rejected enzymes form a gel layer on the membrane surface (Sakaki et al., 2001; Trusek-Holownia and Noworyta, 2007). The stability of this dynamic layer can be improved by crosslinking the enzyme molecules with glutaraldehyde (Wang et al., 2008). It is important to note that there is no covalent binding Gelification on the membrane
Entrapment
Within the Retention in polymeric matrix membrane pores
Attachment on the membrane Physical adsorption
Risk of leaching Diffusional limitation
Covalent binding
Expensive and irreversible binding
Fig. 11.4 Different types of active membrane preparation and their main drawbacks.
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Membrane bioreactors and the production of food ingredients 321 between the support material and the enzymes; the membrane regeneration is thus possible by removing the enzyme gel through backwash with a highpressure water or gas. As physical entrapment in the porous support, this immobilization method is really simple and generally leads to high enzyme loading. However, in both cases all the biocatalyst molecules are not active owing to diffusion limitation or steric obstruction that hinders access of the substrate to the catalytic site of enzyme. It is not obvious to conclude which one of these two methods is the most attractive. Wang et al. (2008), who compared both immobilization methods with the same membrane reactor, observed that the performance is enhanced when the enzymes are on the surface whereas Hilal et al. (2006) reported opposite findings. Finally, enzymes can be attached by noncovalent binding (adsorption through hydrophobic or ionic interactions) or covalent binding on membrane surface. Enzyme adsorption is certainly the simplest and cheapest immobilization method because it can be achieved in one step (by immersion in enzyme solution or filtration) without the use of any activator. Even if protein adsorption is higher on hydrophobic membrane as it is generally reported and demonstrated (Shamel et al., 2007), this method can be applied to various membrane materials from hydrophilic to hydrophobic ones. Recently, Engel et al. (2008) used a strongly basic anion-exchanger membrane for immobilization of b-galactosidase in order to synthesize galacto-oligosaccharides. However, owing to the weakness of the binding force the risk of desorption is high, thus reducing the potential of such active membranes. Nevertheless, according to Tischer and Kasche (1999), immobilization via adsorption method is particularly appropriate when the membrane is used in nonaqueous solvents in which desorption phenomena may be overcome owing to the low solubility of enzymes in such solvents. Compared with noncovalent binding, covalent attachment provides virtually irreversible binding between the amino or carboxyl groups of the enzyme and functional groups of the membrane. It thus avoids biocatalyst leaching and increases enzyme stability especially in nonaqueous reaction media. However, the irreversibility of the linkage may be a serious drawback when the biocatalyst becomes inactive; both biocatalysts and support are unusable. In addition, compared with free enzymes a decrease in enzymatic activity is observed if the catalytic site is involved in the bonding reaction with the support or if the enzyme molecules are immobilized in an inactive conformation. Covalent attachment is generally considered as an expensive method of immobilization. It needs the use of chemicals (i.e. reactive groups such as carbodi-imide, cyanogen bromide, diazonium salts or reagents such like epichlorhydrin or glutaraldehyde) to form the covalent bonds and previous treatments for membrane surface activation (i.e. irradiation or chemical treatments) especially in the case of inorganic supports (Sousa et al., 2001). To overcome some of these drawbacks a new method for enzyme immobilization onto ceramic membranes has been developed at the European
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322 Separation, extraction and concentration processes Membrane Institute (IEM) (Belleville et al., 2001, Magnan et al., 2004). In order to generate functional groups on the membrane surface, the ceramic support is first coated by filtration with polymers (such as gelatine or polyethyleneimine). Then the enzymes are covalently linked to the polymeric layer by glutaraldehyde bonds. Such active membranes are of interest for several reasons linked to their inorganic nature. They can be used under pressure and with a wide range of solvents in particular in supercritical media (Gumi et al., 2007; Lozano et al., 2004). In addition, the ceramic support can be easily regenerated when the enzymes are deactivated. Furthermore, the hydrophilic nature of the polymer coating offers a good environment to preserve enzymes from deactivation by anhydrous conditions. In summary, there are many routes to prepare active membranes and as all of them have advantages and drawbacks, it is not possible to determine the best method. The choice has to be done on a case-by-case basis depending on the type of biocatalyst, membrane material and reactor configuration.
11.3 Applications of membrane bioreactors in food industries During the past two decades, the MBR market worldwide experienced a huge increase from around US$ 1 million in 1990 to US$ 296 million in 2008. It is expected to grow at an annual growth rate of 10.5% to rise to nearly US$ 0.5 billion by 2013 (Anon., 2008). However, these figures concern mainly the market of wastewater treatment driven mainly by legislation and water stress. The industrial applications of MBRs in the food area are much less widespread. Giorno (2008) who analyzed the development of MBRs in terms of publication of patents from 2004 to 2008 reported that patents in food, biotechnology and pharmaceuticals represent only 12% of the total patents whereas water treatments represent 77%. For food applications, the uses of MBRs are mainly involved in the processing of food and beverages (e.g. wine, fruit juices and milk) on the one hand and on the other hand for the production of a wide range of food ingredients obtained by biocatalysis processes (e.g. sugars, organic acids, peptides, esters). 11.3.1 Processing of food and beverages The idea of producing ethanol in a MBR is not really new. In the early 1980s, Cheryan and Mehaia reported the production of ethanol by Saccharomyces cerevisae in a MBR (Cheryan and Mehaia, 1984). They obtained MBR conversions which were 30 times higher than the one observed in batch reactor. In 1991, they associated two MBRs in series in order to produce vinegar from date juice (Mehaia and Cheryan, 1991). The sugars were first converted to ethanol by S. cerevisae in the first MBR. The permeate obtained
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Membrane bioreactors and the production of food ingredients 323 was then used to feed the second MBR in which ethanol was converted into acetic acid by Acetobacter aceti. This two-stage fermentation process exhibited a yield of about 0.5 g acetic acid per gram sugars which is about 75% of the theoretical value and close to that obtained in batch fermentations in current commercial vinegar production. In addition to its high productivity rate, the process allowed the production of a cell-free, clear product stream requiring little or no further downstream processing. More recently, Takaya et al. (2002) examined yeast growth and ethanol fermentation in MBR using a grape juice. They suggested the use of a double vessel membrane reactor, a continuous stirred tank reactor (CSTR) associated with a CRMBR, with which they observed a productivity of dry wine 28 times higher than that in the batch fermentation. Zhang and Lovitt (2006) suggested the use of MBR as a strategy for enhancing malolactic fermentation in wine and cider maturation. Concerning fruit or vegetable juices production, ultrafiltration and microfiltration represent attractive alternatives to conventional processes for juice clarification since they can be carried out in continuous mode without the need of fining agents (Álvarez et al., 2000). However, in order to improve permeate flux and reduce membrane fouling, a pretreatment of juices and pulps with pectinases is usually recommended. In general; this treatment is realized in a tank before juice filtration; however, when this reaction is carried out in a FEMBR, the product inhibition is avoided and the reactor productivity is enhanced compared with a batch system (Bélafi-Bakó et al., 2007). Rodriguez-Nogales et al. (2008), who studied the stability of such reactors for a long term experiment (15 days), observed a significant viscosity reduction (88% below the initial value) and concluded that both operations (pectin hydrolysis and juice clarification) can be achieved in a one step operation with FEMBR. In the dairy industry, the main application of MBRs and more particularly of EMBRs is lactose hydrolysis in order to improve milk digestibility. Grano et al. (2004) investigated the potentiality of a nonisothermal IEMBR. In this reactor, the feed (skimmed milk) flows along both sides of the active membrane prepared with b-galactosidase. Because of the temperature difference applied across the membrane, a thermodialysis flux occurs, from the warm to the cold side of the bioreactor and the lactose transformation occurs during the crossing. However, even if the reactor performance in terms of percentage reduction of the production time is promising and similar to those of batch and fluidized-bed reactors previously reported in the literature (Roy and Gupta, 2003), extrapolating from the laboratory-scale reactor to an industrial-scale process is not straightforward. Another type of reactor is proposed by Novalin et al. (2005). In this instance, the biocatalyst (i.e. b-galactosidase MaxilactTM, Gist-brocades) is retained in the shell side of a hollow-fibre module. Pasteurized skim milk is pumped through the hollow-fibre module (lumen side) and enzymatic solution is circulated inside the shell side. Owing to the diffusion gradient, lactose crosses the membrane
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324 Separation, extraction and concentration processes to the shell side where the reaction takes place. This diffusion reactor shows promising performances (a conversion rate around 80% within laboratoryscale operating conditions) and needs to be investigated at industrial scale. 11.3.2 Production of food ingredients The main applications of enzymatic and whole cell MBRs for food ingredients production are reported in Table 11.1 and 11.2, respectively. First, enzymatic MBRs are much more widely investigated for food ingredients production over the past 10 years than whole cell MBRs. This may be related to the difficulty of using biological cells at an industrial level, particularly owing to a decrease in activity. That may result from a nutriments limitation, product toxicity or mechanical stress, and membrane fouling which induces lower productivity. In addition, during microbial growth, many metabolites or byproducts are released in the fermentation broth at the same time as the product of interest. It is then necessary to consider further stages of separation. Nevertheless, for reactions that implicate several successive enzymes or enzyme cofactors, whole cell MBRs are more attractive than enzymatic ones. As shown in Table 11.2, whole cell MBRs are mostly investigated for the production of organic acids, or small sugars used as sweeteners. Among the various applications reported, lactic acid production is the most studied over the past few years; this is probably because of its possible use as a monomer material for the elaboration of environmentally friendly packaging material, rather than its use as a food additive. However, biosynthesis of lactic acid is an attractive process because it allows the production of high value-added products from lactose, a byproduct of cheese manufacturing. As lactic acid fermentation is characterized by product inhibition, which affects cell growth and metabolism and thus limits the production, the use of MBR is an interesting alternative. Various configurations of MBRs were investigated. Whole cell CRMBRs are the most studied despite the risk of cell denaturation owing to shear stresses. To avoid cell damage, Giorno et al. (2002) suggested working at low axial velocity (laminar flow condition); but in that instance the reactor performance is affected by fouling, which increases the hydraulic resistance of the membrane. To improve filtration performance it was suggested to use iMBR (Kamoshita et al., 1998; Schiraldi et al., 2003). Kwon et al. (2001), who considered that the main drawback of MBRs was low product concentration, suggested connecting two whole cell CRMBRs in series. They reported that 92 g L–1 lactic acid could be produced with a volumetric productivity of 57 g L–1 h–1, and suggested that this attractive value results from the continuous feeding of the second MBR with the permeate and fresh cells from the first reactor. These fresh cells are more resistant to product inhibition. To limit product inhibition it is also possible to connect the MBR to another separation process, for instance to an electrodialysis unit as suggested by Meynial-Salle et al. (2008). Nevertheless, an important disadvantage of MBRs relates to
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Table 11.1 Applications of enzymatic MBRs in agro-food industries Reaction
Objectives
Starch, Starch hydrolysis sweeteners, oligosaccharides
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Production of glucose or maltose syrups, production of maltodextrins Starch hydrolysis and Production of cyclodextrine synthesis cyclodextrins Pectin hydrolysis
Sucrose hydrolysis
Synthesis of poly-b(2-1)-fructan from sucrose Synthesis of galactooligosaccharides (GOS) from lactose
Reactor
References
CRMBR with free enzymes (amylase and/or amyloglucosidase)
Gaouar et al. (1997), Grzeskowiak-Prywecka and Słomińska (2007), Kedziora et al. (2006), Paolucci-jeanjean et al. (2000), Sarbatly and England (2004) Sakinah et al. (2008), Słomińska et al. (2002)
CRMBR with free enzymes [cyclodextrin glucosyl transferase (CGTase)] Production of pectic CRMBR with free oligosaccharides or enzymes (pectinase and/or galacturonic acids polygalacturonase) IEMBR (pectolytic enzyme preparation) Production of high CRMBR with immobilized fructose syrup enzymes (invertase adsorbed on beads) Production of IEMBR fructooligosaccharides (b-fructofuranosidase or fructosyl transferase) Production of prebiotics (GOS)
Belafi Bakó et al. (2007), Kiss et al. (2009), Olano-Martin et al. (2001) Lozano et al. (1990) Tomotani and Vitolo (2007) Hicke et al. (1999), (2006), Nishizawa et al. (2000)
CRMBR with free Czermak et al. (2004), Foda and Lopez-Leiva enzymes (b-galactosidase) (2000), Matella et al. (2006) IEMBR (b-galactosidase) Engel et al. (2008)
Membrane bioreactors and the production of food ingredients 325
Food areas
Food areas
Reaction
Objectives
Reactor
References
Proteins and peptides
Hydrolysis of whey protein
Solubilization of proteins, production of peptides with low allergenicity/or functional properties
CRMBR with free proteases (protease preparation, pepsin, chymotrypsin, carboxypeptidase, alcalase)
Cabrera-Padilla et al. (2009), Cheison et al. (2006a, 2006b), Cheison et al. (2007), Guadix et al. (2006), Mišún et al. (2008), Perea and Ugalde 1996)
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Hydrolysis of other type of proteins (casein, blood, alfafa proteins) Oil and esters
Kapel et al. (2006), Prevot-D’Alvise et al. (2004), Trusek-Holownia (2008), Wei and Chiang (2009)
Alcoholysis, acidolysis Oil engineering, esters CRMBR with free or or transesterification production immobilized lipases Monophasic IEMBR (lipase) Oil hydrolysis
Production of fatty acids, monoglycerides, diglycerides
Xu et al. (2000), Trusek-Holownia and Noworyta (2007) Giorno et al. (1997), Gumi et al. (2007), Hernández et al. (2006), Lozano et al. (2002), (2004), Pomier et al. (2005), Trusek-Holownia and Noworyta (2007) CRMBR with free lipases Gan et al. (1998) Monophasic IEMBR (lipase)
Garcia et al. (1992)
Biphasic IEMBR (lipase)
Giorno et al. (1997), Goto et al. (1992), Knezevic et al. (2004), Merçon et al. (2000), Pugazhenthi and Kumar (2004), Sachan et al. (2006), Shamel et al. (2007), Tan et al. (2002), Wang et al. (2008)
326 Separation, extraction and concentration processes
Table 11.1 Continued
Membrane bioreactors and the production of food ingredients 327 Table 11.2 Applications of whole cell MBRs in agro-food industries Objectives
Reactor
References
Production of lactic acid
CRMBR with free cells (Lactobacillus rhamnosus; Lactobacillus bulgaricus) Filtration unit spiral-sheet polymeric membrane with adsorbed cells (Bifidobacterium longum and B. helveticus) associated to a CSTR iMBR with free cells (Lactococcus lactis; Lactobacillus delbruekii)
Kwon et al. (2001), Giorno et al. (2002) Shahbazi et al. (2005)
Production of xylitol
CRMBR with free cells (Candida spp.)
Ko et al. (2008), Kwon et al. (2006)
Production of palatinose
Hollow-fibre membrane reactor with free cells immobilized in shell side (Serratia plymuthica)
Krastanov et al. (2007)
Kamoshita et al. (1998), Schiraldi et al. (2003)
Production of CRMBR with free cells (Azotobacter Cheze-Lange et al. (2002) microbial alginate vinelandii) Production of manitol
CRMBR with free cells (Leuconostoc von Weymarn et al. mesenteroides) (2002)
Production of succinic acid
CRMBR with free cells (Anaerobiospirillum succiniciproducens) associated with an electrodialysis unit
Meynial et al. (2008)
the production of metabolite or by-products with low interest. Hence, the recovery of the desired product needs further separating steps. With regard to food applications of enzymatic membrane reactors, three main areas can be distinguished: (1) starch-derived-products, sweeteners and oligosaccarides production, (2) protein hydrolysates production, (3) fatty-acid esters production and oil modification. Starch conversion into smaller assimilable sugars is the most studied reaction (Gaouar et al., 1997; Grześkowiak-Pryweck and Słomińska, 2007; Kedziora et al., 2006; Paolucci-Jeanjean et al., 2000; Sarbatly and England, 2004). Usually this reaction is carried out in a FEMBR using simultaneously amylolytic enzymes and debranching enzymes (Gaouar et al., 1997) or using liquefied starch (i.e. maltodextrin) as substrate (Grześkowiak-Pryweck and Słomińska, 2007). This was intended to produce an increased yield while preventing accumulation of limit b-dextrins in the reactor, these molecules being suspected of fouling the membrane. Paolucci-Jeanjean et al. (2000) studied the degradation of raw cassava starch into maltodextrin using Termamyl™, a thermostable a-amylase supplied by NOVO, and showed that the major drawbacks observed were the membrane fouling owing to accumulation of high molecular weight products on a one hand, and enzyme inactivation on © Woodhead Publishing Limited, 2010
328 Separation, extraction and concentration processes the other hand. In order to achieve a high level of conversion and to improve the control of system performance, Paolucci-Jeanjean et al. suggested filling the reactor with a pre-treated solution before starting the continuous feeding of the closed-loop membrane reactor with the raw starch solution. However, in a more recent critical review, Sarbatly and England (2004) presented different potential opportunities to prevent membrane fouling such as the use of particles as turbulence promoters or the use of an active membrane. In the latter case, the high molecular weight products which accumulate at the membrane surface are hydrolyzed and the degraded products can flow through the pore. Starch can also be used as a substrate for cyclodextrin production. Słomińska et al. (2002) reported that compared with the batch process, the use of MBR with cyclodextrin glucosyl transferase (CGTase) increases the process efficiency in particular when the starch concentration is high (i.e., 20% w/v). Sakinah et al. (2008) who studied the same reaction, showed that the use of gelatinized starch is unfavourable even if the enzyme tested showed a higher affinity for this substrate. Actually, the use of gelatinized starch leads to higher membrane fouling owing to the cake deposition associated with large swollen tapioca starch molecules. As reported in Table 11.1, the production of oligosaccharides can be achieved in EMBR either by polysaccharide hydrolysis (i.e. hydrolysis of pectin), or by a synthesis reaction starting from sucrose (synthesis of fructo-oligosaccharides) or lactose (synthesis of galacto-oligosaccharides). However, for fructo-oligosaccharides, it is more convenient to produce these molecules from hydrolysis of fructan-containing crop materials or directly from agave juice which is known as a natural source of fructo-oligosaccharides (Ortiz-Basurto et al., 2008). The use of enzymatic FEMBRs is also attractive to produce protein hydrolysates in particular from whey hydrolysis (Cabrera-Padilla et al., 2009; Cheison et al., 2006 a,b; Cheison et al., 2007; Guadix et al., 2006; Mišún et al., 2008; Perea and Ugalde, 1996) or from other protein sources (Kapel et al., 2006; Prevot-D’Alvise et al., 2004; Trusek-Holownia, 2008; Wei and Chiang, 2009); this is an attractive way to produce various functional and bioactive peptides. Enzymatic hydrolysis of whey proteins has been demonstrated to be an excellent method to reduce their allergenicity and the peptides obtained are widely used as food ingredients in energy-providing drinks, hypoallergenic formulae and enteral diets for children and sick adults. Although hydrolysis of protein is generally carried out in a batch reactor, this reaction can be achieved with a CRMBR. Cheison et al. (2006a, 2006b) who studied the hydrolysis of whey with protease N ‘Amano’ (Amano Enzymes Co., Nagoya, Japan) in FEMBR underlined the positive role of retentate temperature which influences the solubility of protein and the whey viscosity thus leading to higher permeate fluxes. According to them, the reactor performances depend on the flux values and enzyme concentration. At low permeate flux as well as low enzyme concentration, enzyme activity is limited
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Membrane bioreactors and the production of food ingredients 329 owing to substrate inhibition whereas high permeate flux leads to enzyme leakage. Thus, hydrodynamic properties are crucial for reactor robustness. Cabrera-Padilla et al. (2009) investigated a novel enzymatic MBR for the production of whey hydrolysates with a low content of phenylalanine. Instead of free enzymes, they chose to use carboxypeptidase A (CPA), immobilized on agarose gel particles, thus avoiding enzyme autolysis, an important drawback of processes involving proteases, and preserving the biocatalyst activity for several runs. In addition, the reactor was continuously fed with whey previously hydrolyzed with immobilized chymotrypsin to a 15.3% degree of hydrolysis. Under these conditions, they observed that their EMBR had a better performance than a conventional batch reactor associated with a diafiltration unit. Finally, the use of enzymatic MBR for the biotransformation of oils and fats has been widely reported. These bioreactors generally involve lipases or esterases which can catalyze a wide range of reactions such as hydrolysis, alcoholysis, transesterifications, aminolysis and enantiomer resolution (Hasan et al. 2006; Jaeger and Eggert, 2002). In the food industry, the major application of EMBRs is esters synthesis for the production of emulsifiers and aroma compounds and, to a lesser extent, oil hydrolysis for production of free fatty acids (FFA) and mono or diglycerides as reported in Table 11.1. Compared with the conventional process for fats and oil hydrolysis, the NaOH-catalyzed hydrolysis process, which requires high pressure of about 4.82 MPa (or more) and high temperature of about 250 °C, and even though the chemical process gives high conversion rate (97–98%), the biocatalytic route is really more attractive since it can be carried out at ambient temperature and pressure with a significant decrease of waste (Wang et al., 2008). In a first work, Garcia et al. (1992) studied the hydrolysis of milkfat in an IEMBR; the enzymes were adsorbed onto polypropylene microporous membrane and the milkfat was partially hydrolyzed during the membrane crossing. However, such a reactor leads to a mixture of free fatty acids and partially hydrolyzed milkfat. In 1998, Gan et al. (1998) investigated another type of reactor. They associated a stirred tank reactor with a de-emulsifier; the reactor previously filled with substrate (sunflower oil) and lipase was continuously fed with an aqueous buffer. The resulting emulsion was pumped through the de-emulsifier and whereas the oil phase was directly recycled in the reactor, the water phase was sent to an ultrafiltration unit. The biocatalyst, which was concentrated in the retentate, was recycled whereas glycerol was removed with the permeate. Gan et al. (1998) expected to increase the recovery of FFAs by a continuous removal of glycerol; but they observed only marginal improvements in overall reaction yield. They supposed that this was because of the unsuccessful separation of the free fatty acids produced from both the de-emulsified aqueous and oil phases. To overcome this drawback and obtain FFAs separately, most studies (Giorno et al., 1997; Goto et al., 1992; Knezevic et al., 2004; Merçon et al., 2000; Pugazhenthi and Kumar, 2004; Sachan et al., 2006; Shamel et al., 2007; Tan et al., 2002; Wang et al., 2008) refer to the use of biphasic IEMBR,
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330 Separation, extraction and concentration processes with the enzymatic membrane placed between two immiscible phases, the oil and an aqueous buffer, where the fatty acids were released. However, owing to high mass transfer resistance, the productivity of such biphasic IEMBRs is generally very low. They are thus limited to systems working in a reaction limited regime. For ester synthesis, because these reactions generally involve low molecular weight substrates having low water solubility, the use of IEMBRs with nonaqueous solvent seems to be an attractive alternative for ester bioproduction. Lozano et al. (2002) and Magnan et al. (2004) have successfully achieved the synthesis in organic solvent of butyl, butyrate and butyl laurate, respectively, using a monophasic IEMBR with a hybrid organic–inorganic catalytic membrane developed at IEM and briefly presented above. These results confirm that the hydrophilic behaviour of the inert protein ensures a proper environment for the enzymes, which are preserved from inactivation in nonaqueous media. The immobilized enzyme, (Candida antarctica lipase B (CALB)), exhibited a particularly high stability, a half-life of 202 days, (Lozano et al., 2002). Owing to the drawbacks of organic solvents and above all their toxicity, some studies used more environmentally friendly solvents such as supercritical carbon dioxide (SC CO2). Despites the numerous advantages of SC CO2, among which are the tunability of solvent properties and simple downstream processing features, only a few examples of EMBRs using SC CO2 as a reaction medium have been reported. Lozano et al. (2004) compared the butyl butyrate synthesis catalyzed by CALB in IEMBR using either SC CO2 or organic solvents (acetonitrile, acetone and hexane) as reaction media. Both activity and selectivity were enhanced in supercritical medium. However, it must be noted that the IEMBR described here was used in a batch mode without membrane permeation; the interest of the system is thus limited. Nevertheless, when the reaction is carried out in continuous dead-end filtration mode, the production of ester is enhanced compared with experiments carried out in batch and semi-batch modes as shown by Gumi et al. (2007) for butyl laurate synthesis in SC CO2. This is because the convective transfer of substrate through the membrane is much more efficient than the diffusive one. Finally, the use of EMBR for structured lipids or engineered oil production was studied by Xu et al. (2000) for the reaction between medium-chain triacylglycerols (MCT) and n-3 polyunsaturated fatty acids from fish oil using Lipozyme™ (Novo Nordisk) as a biocatalyst in a membrane reactor. They reported that the percentage incorporation of polyunsaturated fatty acids into MCT was increased by about 15% over 80 h by simultaneous separation of the released medium-chain fatty acids compared with a batch experiment. More recently, Pomier et al. (2005) investigated the enzymatic modification of castor oil through the interesterification of castor oil triglycerides and methyl oleate. The reaction catalyzed by CALB was carried out in an IEMBR and SC CO2 was used as thinning agent in order to strongly decrease oil viscosity and thus to enable membrane filtration. The stability of the immobilized
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Membrane bioreactors and the production of food ingredients 331 enzymes was checked during a 25 h period and a conversion around 30% was observed. It is worth noting that the concentration of methyl ricinoleate was higher in the permeate than in the retentate showing that reaction happened mainly in the pores of the membrane where the contact between enzymes and substrates was much more favourable than at the internal surface.
11.4 Future trends Throughout this review, the potential of MBR for food processing or food ingredients production is clearly shown. However, further investigations are needed in order to guarantee the success of these reactors at the industrial level, particularly regarding their performance, integration and potential for process intensification. Further study is required on: (1) membrane materials, (2) biocatalyst engineering, (3) process engineering. Firstly, as suggested by Meng et al. (2009) in a recent review, a comprehensive investigation should be performed to understand, control and reduce membrane fouling, and particularly to avoid severe fouling. This might lead to the development of new membrane materials or the modification of those already existing in order to obtain high transfer rates. This is particularly important to improve MBR productivity particularly for whole cell MBR. The second challenge is related to research in genetic and enzyme engineering. Investigations are required to produce at lower price enzymes showing higher activity, specificity and stability in order to increase economical interest of EMBR. Last but not least, developing new ideas and concepts as well as specific modelling tools to enable easier process integration and scale changes appears today as an holistic way to optimize the performance of MBRs and to enlarge their industrial applications. Indeed, as stressed by Drews and Kraume (2005), MBRs should be considered as hybrid reactors and not as the juxtaposition of independent bioreactor and membrane filtration units for the optimization, the modelling and the scaling-up of such processes.
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332 Separation, extraction and concentration processes Anon. (2008), ‘Membrane market is experiencing a strong growth’, Membr Technol, October 2008, 10–11. Bartling K, Thompson J U S, Pfromm P H, Czermak P and Rezac M E (2001), ‘Lipasecatalysed synthesis of geranyl acetate in n-hexane with membrane-mediated water removal’, Biotechnol Bioeng, 75, 676–681. Bélafi-Bakó K, Eszterle M, Kiss K, Nemestóthy N and Gubicza L (2007), ‘Hydrolysis of pectin by Aspergillus niger polygalacturonase in a membrane bioreactor’, J Food Eng, 78, 438–442. Belfort G (1989), ‘Membranes and bioreactors: a technical challenge in biotechnology’, Biotechnol Bioeng, 33, 1047–66. Belleville M-P, Lozano P, Iborra J L and Rios G M (2001), ‘Preparation of hybrid membranes for enzymatic reaction’, Sep Purif Technol, 25, 229–233. Bódalo A, Gómez J L, Gómez E, Máximo M F and Montiel M C (2004), ‘Study of l-aminoacylase deactivation in an ultrafiltration membrane reactor’, Enzyme Microb Technol, 35, 261–266. Cabrera-Padilla R Y, Pinto G A, Giordano R L C and Giordano R C (2009), ‘A new conception of enzymatic membrane reactor for the production of whey hydrolysates with low contents of phenylalanine’, Process Biochem, 44, 269–276. Cheison S C, Wang Z and Xu S-Y (2006a), ‘Hydrolysis of whey protein isolate in a tangential flow filter membrane reactor. I. Characterisation of permeate flux and product recovery by multivariate data analysis’, J Membr Sci, 283, 45–56. Cheison S C, Wang Z and Xu S-Y (2006b), ‘Hydrolysis of whey protein isolate in a tangential flow filter membrane reactor. II. Characterisation for the fate of the enzyme by multivariate data analysis’, J Membr Sci, 286, 322–332. Cheison S C, Wang Z and Xu S-Y (2007), ‘Use of response surface methodology to optimise the hydrolysis of whey protein isolate in a tangential flow filter membrane reactor’, J Food Eng, 80, 1134–1145. Cheryan M and Mehaia M A (1984), ‘Ethanol production in a membrane recycle bioreactor – conversion of glucose using Saccharomyces cerevisiae’, Process Biochem, 19, 204–208. Cheryan M and Mehaia M A (1986), ‘Membrane bioreactors’, In: W C McGregor, Membrane separation in biotechnology, New York, Marcel Dekker, 255–301. Chèze-Lange H, Beunard D, Dhulster P, Guillochon D, Cazé A M, Morcellet M, Saude N and Junter G-A (2002), ‘Production of microbial alginate in a membrane bioreactor’, Enzyme Microb Technol, 30, 656–661. Czermak P, Ebrahimi M, Grau K, Netz S, Sawatzki G and Pfromm P H (2004), ‘Membrane-assisted enzymatic production of galactosyl-oligosaccharides from lactose in a continuous process’, J Membr Sci, 232, 85–91. Del Amor Villa E M and Wichmann R (2005), ‘Membranes in enzymatic synthesis of biotensides from renewable sources’, Catal Today, 104, 318–322. Drews A and Kraume M (2005), ‘Process improvement by application of membrane bioreactors’, Chem Eng Res Des, 83(A3), 276–284. Engel L, Ebrahimi M and Czermak P (2008), ‘Membrane chromatography reactor system for the continuous synthesis of galactosyl-oligosaccharides’, Desalination, 224, 46–51. Foda M I and Lopez-Leiva M (2000), ‘Continuous production of oligosaccharides from whey using a membrane reactor’, Process Biochem, 35, 581–587. Gan Q, Rahmat H and Weatherley L R (1998), ‘Simultaneous reaction and separation in enzymatic hydrolysis of high oleate sunflower oil – evaluation of ultrafiltration performance and process synergy’, Chem Eng J, 71, 87–96. Gaouar O, Aymard C, Zakhia N and Rios G M (1997), ‘Enzymatic hydrolysis of cassava starch into maltose syrup in a continuous membrane reactor’, J Chem Technol Biotechnol, 69, 367–375. Garcia H S, Malcata F X, Hill C G and Amundson C H (1992), ‘Use of Candida rugosa
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334 Separation, extraction and concentration processes M, Guillochon D and Dhulster P (2006), ‘Production, in continuous enzymatic membrane reactor, of an anti-hypertensive hydrolysate from an industrial alfalfa white protein concentrate exhibiting ACE inhibitory and opioid activities’, Food Chem, 98, 120–126. Kedziora P, Le Thanh J, Lewandowicz G and Prochaska K (2006) ‘An attempt to application of continuous recycle membrane reactor for hydrolysis of oxidised derivatives of potato starch’, J Membr Sci, 282, 14–20. Kiss K, Nemestóthy N, Gubicza L and Bélafi-Bakó K (2009), ‘Vacuum assisted membrane bioreactor for enzymatic hydrolysis of pectin from various agro-wastes’, Desalination, 241, 29–33. Knezevic Z, Kukic G, Vukovic M, Bugarski B and Obradovic B (2004), ‘Operating regime of a biphasic oil/aqueous hollow-fibre reactor with immobilized lipase for oil hydrolysis’, Process Biochem, 39, 1377–1385. Ko C-H, Chiu P-C, Yang C-L and Chang K-H (2008), ‘Xylitol conversion by fermentation using five yeast strains and polyelectrolyte-assisted ultrafiltration’, Biotechnol Lett, 30, 81–86. Krastanov A, Blazheva D and Stanchev V (2007), ‘Sucrose conversion into palatinose with immobilized Serratia plymuthica cells in a hollow-fibre bioreactor’, Process Biochem, 42, 1655–1659. Kwon S, Yoo I-K, Lee W G, Chang H N and Chang Y K (2001), ‘High-rate continuous production of lactic acid by Lactobacillus rhamnosus in a two-stage membrane cellrecycle bioreactor’, Biotechnol Bioeng, 73, 25–34. Kwon S-G, Park S-W and Oh D-K (2006), ‘Increase of xylitol productivity by cellrecycle fermentation of Candida tropicalis using submerged membrane bioreactor’, J Biosci Bioeng, 101, 13–18. Long W S, Kow P C, Kamaruddin A H and Bhatia S (2005), ‘Comparison of kinetic resolution between two racemic ibuprofen esters in an enzymatic membrane reactor’, Process Biochem, 40, 2417–2425. Lozano P, Manjón A, Iborra J L, Cánovas M and Romojaro F (1990), ‘Kinetic and operational study of a cross-flow reactor with immobilized pectolytic enzymes’, Enzyme Microb Technol, 12, 499–505. Lozano P, Pérez-Marín A B, De Diego T, Gómez D, Paolucci-Jeanjean D, Belleville M-P, Rios G M and Iborra J L (2002), ‘Active membranes coated with Candida antarctica lipase B: preparation and application for continuous butyl butyrate synthesis in organic media’, J Membr Sci, 201, 55–64. Lozano P, Víllora G, Gómez D, Gayo A B, Sánchez-Conesa J A, Rubio M and Iborra J L (2004), ‘Membrane reactor with immobilized Candida antarctica lipase B for ester synthesis in supercritical carbon dioxide’, J. Supercrit Fluids, 29, 121–128. Magnan E, Catarino I, Paolucci-Jeanjean D, Preziosi-Belloy L and Belleville M-P (2004), ‘Immobilisation of lipase on a ceramic membrane: activity and stability’, J Membr Sci, 241, 161–166. Matella N J, Dolan, K D and Lee Y S (2006), Comparison of galactooligosaccharide production in free-enzyme ultrafiltration and in immobilized-enzyme systems’, J Food Sci, 71, C363–C368. Mateo C, Palomo J M, Fernandez-Lorente G, Guisan J M and Fernandez-Lafuente R (2007), ‘Improvement of enzyme activity, stability and selectivity via immobilization techniques’, Enzyme Microb Technol, 40, 1451–1463. Mehaia M A and Cheryan M (1991), ‘Fermentation of date extracts to ethanol and vinegar in batch and continuous membrane reactors’, Enzyme Microb Technol, 13, 257–261. Meng F, Chae S R, Drews A, Kraume M, Shin H-S and Yang F (2009), ‘Recent advances in membrane bioreactors (MBRs): membrane fouling and membrane material’, Water Res, 43, 1489–1512. Merçon F, Sant’Anna G L and Nobrega R (2000), ‘Enzyme hydrolysis of babassu oil in a membrane reactor’, JAOCS, 77(10), 1043–1048. © Woodhead Publishing Limited, 2010
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336 Separation, extraction and concentration processes (2008), ‘Cyclodextrin production in hollow fiber membrane reactor system: effect of substrate preparation’, Sep Purif Technol, 63, 163–171. Sanchez Marcano J and Tsotsis T T (2002), ‘Membrane bioreactors’ in Catalytic Membranes and Membrane Reactors, Weinheim, Wiley VCH, 133–168. Sarbatly R and England R (2004), ‘Critical review of membrane bioreactor system used for continuous production of hydrolyzed starch’ Chem Biochem Eng Q, 18(2), 155–165. Schiraldi C, Adduci V, Valli V, Maresca C, Giuliano M, Lamberti M, Cartenì M and De Rosa M (2003), ‘High cell density cultivation of probiotics and lactic acid production’, Biotechnol Bioeng, 82, 213–222. Shahbazi A, Mims M R, Li Y, Shirley V, Ibrahim S A and Morris A (2005), ‘Lactic acid production from cheese whey by immobilized bacteria’, Appl Biochem Biotechnol, 121–124, 529–540. Shamel M M, Ramachandran K B, Hasan M and Al-Zuhair S (2007), ‘Hydrolysis of palm and olive oils by immobilised lipase using hollow fibre reactor’, Biochem Eng J, 34, 228–235. Słomińska L, Szostek A and Grześkowiak A (2002), ‘Studies on enzymatic continuous production of cyclodextrins in an ultrafiltration membrane bioreactor’, Carbohydr Polym, 50, 423–428. Sousa H A, Rodrigues C, Klein E, Afonso C A M and Crespo J G (2001), ‘Immobilisation of pig liver esterase in hollow fibre membranes’, Enzyme Microb Technol, 29, 625–634. Takaya M, Matsumoto N and Yanase N (2002), ‘Characterization of membrane bioreactor for dry wine production’, J Biosci Bioeng, 93(2), 240–244. Tan T, Wang F and Zhang H (2002), ‘Preparation of PVA/chitosan lipase membrane reactor and its application in synthesis of monoglyceride’, J Mol Catal B: Enzym, 18, 325–331. Tischer W and Kasche V (1999), ‘Immobilized enzymes: crystals or carriers?’ Trends biotechnol, 17, 326–335. Tomotani E J and Vitolo M (2007), ‘Production of high-fructose syrup using immobilized invertase in a membrane reactor’, J Food Eng, 80, 662–667. Torras C, Nabarlatz D, Vallot G, Montané D, Garcia-Valls R (2008), ‘Composite polymeric membranes for process intensification: enzymatic hydrolysis of oligodextrans’, Chem Eng J, 144, 259–266. Trusek-Holownia A (2008), ‘Production of protein hydrolysates in an enzymatic membrane reactor’, Biochem Eng J, 39, 221–229. Trusek-Holownia A and Noworyta A (2007), ‘An integrated process: ester synthesis in an enzymatic membrane reactor and water sorption’ J Biotechnol, 130, 47–56. Valadez-Blanco R, Ferreira F C, Ferreira Jorge R and Livingston A G (2008), ‘A membrane bioreactor for biotransformations of hydrophobic molecules using organic solvent nanofiltration (OSN) membranes’, J Membr Sci, 317, 50–64. Vane L M (2005), ‘A review of pervaporation for product recovery from biomass fermentation processes’, J Chem Technol Biotechnol, 80, 603–629. von Weymarn N, Kiviharju K and Leisola M (2002), ‘High-level production of d-mannitol with membrane cell-recycle bioreactor’, J Ind Microbiol Biotechnol, 29, 44–49. Wang Y, Hu Y, Xu J, Luo G and Dai Y (2007), ‘Immobilization of lipase with a special microstructure in composite hydrophilic CA/hydrophobic PTFE membrane for the chiral separation of racemic ibuprofen’, J Membr Sci, 293, 133–141. Wang Y, Xu J, Luo G and Dai Y (2008), ‘Immobilization of lipase by ultrafiltration and cross-linking onto the polysulfone membrane surface’, Bioresour Technol, 99, 2299–2303. Wei J-T and Chiang B-H (2009), ‘Bioactive peptide production by hydrolysis of porcine blood proteins in a continuous enzymatic membrane reactor’, J Sci Food Agric, 89, 372–378.
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Membrane bioreactors and the production of food ingredients 337 Won K, Hong J-K, Kim K J and Moon S-J (2006), ‘Lipase-catalyzed enantioselective esterification of racemic ibuprofen coupled with pervaporation’, Process Biochem, 41, 264–269. Xu J, Wang Y, Hu Y, Luo G and Dai Y (2006), ‘Candida rugosa lipase immobilized by a specially designed microstructure in the PVA/PTFE composite membrane’, J Membr Sci, 281, 410–416. Xu X, Skands A, Jonsson G and Adler-Nissen J (2000), ‘Production of structured lipids by lipase-catalysed interesterification in an ultrafiltration membrane reactor’, Biotechnol Lett, 22, 1667–1671. Zhang D and Lovitt R W (2006), ‘Strategies for enhanced malolactic fermentation in wine and cider maturation’, J Chem Technol Biotechnol, 81, 1130–1140. Ziobrowski Z, Kiss K, Rotkegel A, Nemestóthy N, Krupiczka R and Gubicza L (2009), ‘Pervaporation aided enzymatic production of glycerol monostearate in organic solvents’, Desalination, 241, 212–217.
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338 Separation, extraction and concentration processes
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339
Part II Separation technologies in the processing of particular foods and nutraceuticals
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340 Separation, extraction and concentration processes
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Separation technologies in dairy and egg processing 341
12 Separation technologies in dairy and egg processing G. Gésan-Guiziou, INRA, France
Abstract: A description is given of the separation processes applied in the dairy and egg-processing sectors. The main focus is on the use of membrane and chromatography processes for protein separation, in accordance with the developed applications in these two food sectors, but an overview of all the separation processes implemented throughout the processing chains (from product reception to wastewater treatment) is also given. Characteristics of the food products are presented and the performance of the main separation techniques used on an industrial scale, in particular with respect to the operating modes and properties of the obtained fractions, is discussed. Finally, some of the options that food companies need to consider over the next decade, such as manufacture of isolated peptides and proteins fractions with specific functions and reduction of environmental impacts, are explored. Key words: dairy industry, egg-processing industry, proteins, membrane processes, chromatography.
12.1 Introduction During the past three decades, in parts of the world where food products are readily available, increasing industrial demand has been directed towards large-scale concentration, extraction and purification procedures of food components. Proteins and, to a lesser extent, lipids have been exploited not only for their basic nutritional contribution owing in particular to their essential amino acids, fatty acids and phospholipids composition, but also for their functional and biological properties which make them potential ingredients of health-promoting foods. Functional ingredients have largely been produced for modifying or enhancing the textural and rheological characteristics of © Woodhead Publishing Limited, 2010
342 Separation, extraction and concentration processes food stuffs: for instance they emulsify fat, bind and entrap water, increase the viscosity of liquids, and form gels of different characteristics. Specific food components are also associated with biological activities such as antimicrobial, antihypertensive, anticancer and opioid activities and electrolyte transfer. In the food sector, the dairy industry has undoubtedly developed the most advanced extraction/separation procedures for concentration and fractionation of molecules from milk and its derivatives. Over the last 40 years, membrane processes have become major tools for the separation of dairy components. The first membrane development in the separation procedures of milk components occurred in the late 1960s with the advent of membrane separation and since then a new industry has spawned for whey treatment as well as new avenues for cheesemaking. The separation techniques, and particularly membrane separations, have been implemented throughout the dairy processing chain: milk reception, cheesemaking, serum protein concentration, fractionation of protein, and effluents treatment. Simultaneously, the egg processing industry has seen changes, including high-speed egg-breaking machines, improved pasteurization technology and improved multistage spray driers (Stadelman and Cotteril, 1995). The egg production has sharply increased during the last 20 years (reaching 63 million tonnes of hen eggs in 2007, which corresponds to approximately one thousand billion eggs on a basis of 16 eggs per kg, Nau et al., 2010) and there has been a continuing growth of processed egg products. Today, approximately 30% of the worldwide consumption of eggs is in the form of processed egg products (Froning, 2008), approximately 25% of which is European consumption (Nau et al., 2010). Owing to their multifunctional properties, many of these egg products are used as ingredients in various food applications. But, innovative studies reveal the diversity of chemical properties, structure and function of egg components that has fuelled increasing demand to more fully utilize egg products and has led to the development of new industrial separation processes mainly based on chromatography for the extraction of components (proteins, lipids) from either the egg white or yolk. The aim of this chapter is to give a description of the separation processes applied in the dairy and egg-processing sectors. The main focus is on the use of membrane and chromatography processes for protein separation, in accordance with the developed applications in these two food sectors, but an overview of all the separation techniques implemented throughout the processing chains (from product reception to the wastewater treatment) is also given. The characteristics of the food products are presented and the performance of the main separation techniques used at industrial scale, in particular with respect to the operating modes and properties of the obtained fractions, is discussed. Finally, some of the options that food companies need to consider over the next decade, such as manufacture of isolated peptides and proteins fractions with specific functions and reduction of environmental impacts, are explored.
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Separation technologies in dairy and egg processing 343
12.2 The dairy industry and composition of dairy products It would be impossible to develop a separation process for any group of milk components without prior thorough knowledge of their physicochemical characteristics. 12.2.1 Milk Milk is defined as the secretion of the mammary glands of mammals, its primary natural function being nutrition of the young. In this chapter, the word ‘milk’ will refer to ‘normal’ milk of healthy cows. Milk is a complex fluid, with a constant pH around 6.5–6.7 at ambient temperature. A classification of the principal milk constituents is given in Table 12.1 and Fig. 12.1. The main constituents are lactose (48–50 g L–1), fat (34–44 g L–1), proteins Table 12.1 Approximate composition of cows’ milk and association state Components Concentration (g L–1) Water Fat Lactose Proteins Caseins Soluble proteins Ashes (minerals and salts)
Size (mm) or molecular Association state weight (order of magnitude) (g mol–1 or Da)
870–875 34–44 0.15–15 mm 48–50 342 Da 32–35 25–28 50–500 mm 5–7 14 200–150 000 Da
Solvent Separated phase Solution
8–9
Solution
Ions
Aggregates = micelles Mono-oligomers
Micro-organisms
Soluble proteins Casein micelles
Lactose, small organic molecules
Fat globules
Somatic cells
(mm) 0.0001
0.001
0.01
1
10
100
Microfiltration
Nanofiltration Reverse osmosis
0.1
Ultrafiltration
Fig. 12.1 Approximate particle sizes for which separation by means of membrane filtration can be applied. The size of some milk components is also indicated in comparison with membrane pore size range.
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344 Separation, extraction and concentration processes (32–35 g L–1 expressed as nitrogen, N ¥ 6.38) and minerals (8–9 g L–1), but many other small species, such as urea and vitamins, are present in the soluble phase of milk. Lactose is the distinctive carbohydrate of milk, composed of glucose and galactose. Fat is present in milk in the form of fat globules from 0.15 to 15 mm diameter. There are numerous small fat globules representing a small fraction of the fat but a large surface area, and a few large ones comprising a larger fat percentage (Mulder and Walstra, 1974). These globules are naturally surrounded by their native milk fat globule membrane, which is composed mainly of phospholipids, cholesterol, proteins, and enzymes. Proteins correspond to about 95% of the total nitrogen of the milk. Bovine milk contains numerous proteins which are classically divided into two major groups. First, the caseins, which are insoluble at their isoelectric point (pH = 4.6 at temperature >8 °C), are associated into large globular aggregates, called casein micelles. The composition and structure of the casein micelles have been studied for more than 40 years, and are still not totally elucidated (Horne, 2006; Qi, 2007). They are made of four distinct caseins, as1, as2, b, and k in the proportions 3:1:3:1, and 8% in mass of calcium and phosphate, often called the colloidal phosphate. The structural model accepted most widely has a roughly spherical, core-shell structure, with outer diameters ranging from 50 to 500 nm. The core is generally described as a homogeneous web of caseins in which calcium phosphate nanoclusters are distributed randomly. The shell is essentially made of k-caseins that extend into the aqueous phase as a polyelectrolyte brush producing steric and electrostatic repulsions between micelles. The casein micelles are ‘soft’ and ‘dynamic’ colloids. The second group are the serum proteins, also called soluble proteins because they do not precipitate in the ionic environment of milk when rennet is added or acidification occurs down to pH = 4.6. They are mainly composed of b-lactoglobulin (~ 3.2 g L–1), a-lactalbumin (~ 1.2 g L–1), bovine serum albumin, BSA (~ 0.4 g L–1), immunoglobulins (~ 0.7 g L–1), and minor proteins lactoferrin (LF) and lactoperoxidase (LP). They are largely present in milk in molecular form or as very small oligomers (Table 12.2). The mineral fraction (8–9 g L–1) contains both anions and cations (Table 12.3), which are in dynamic equilibrium among themselves in solution and between solution and proteins (Fig. 12.2). Changing the external conditions of milk, such as pH and temperature, may cause alterations in the mineral equilibriums that could induce modifications in the structure and stability of casein micelles (owing to colloidal phosphate), and then changes in the performance of fractionation processes. Milk collected by the dairy plant also contains a microbial flora formed by numerous species illustrating the contaminations of milk by the udder, the milking machine, the local farm atmosphere, the bulk tank and the transportation equipment. These bacteria account for about 0.01% of the volume of milk of healthy cows. Somatic cells, coming from the bovine © Woodhead Publishing Limited, 2010
Separation technologies in dairy and egg processing 345 Table 12.2 Characteristics (concentration, isoelectric point, molecular weight) and potential biological functions of some proteins found in milk Concentration Isoelectric Molecular Biological (g L–1) point weight (kg functions* mol–1 or kDa) Caseins 25–28 ≈4.6 14–22 as1, as2, k, b
Iron carrier, immunomodulator, precursor for bioactive peptides
b-lactoglobulin 3.2 5.4 18.4 (dimer in milk 36.8)
Retinol carrier, potential antioxidant, precursor for bioactive peptides, binds fatty acids
a-lactalbumin 1.2 4.4 14.2
Lactose synthesis in mammary gland, calcium carrier, immunomodulator, precursor for bioactive peptide
Immunoglobulins 0.5–1.0 5-8 150–1000
Specific immune protection (antibodies and complement system), potential precursor for bioactive peptides
Bovine serum 0.4 5.1 66.3 albumin
Precursors of bioactive peptides
Lactoferrin 0.2 7.9 80
Antimicrobial, antioxidative, anticarcinogenic, anti-inflammatory, iron transport, cell growth, regulation, precursor for bioactive peptides, immunomodulator
Lactoperoxidase 0.03 9.6 78
Antimicrobial, synergetic effect with immunoglobulins and lactoferrin
*Adapted from Korhonen and Pihlanto (2007).
mammary gland (leucocytes, macrophage and epithelial cells) are also normally present in milk (100 000–400 000 cells mL–1). One particular property of milk is that it is not stable. Several changes occur as a result of physical and chemical changes, for instance when temperature
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346 Separation, extraction and concentration processes Table 12.3 Mineral composition of cows’ milk Component
Total concentration in milk (mg kg–1)
Concentration in the soluble phase (mg kg–1)
Calcium, Ca Magnesium, Mg Sodium, Na Potassium, K Chloride, Cl Phosphorus, P Non-colloidal phosphate (expressed in P) Citrate
1250 115 425 1600 1100 950 720
350 70 400 1500 1100 420 300
1650
1500
Non-protein fraction CaCit–
CaHPO4 3b
2b
3a
2a
1a
Ca2+
1b
Colloidal fraction (casein micelles)
HPO42–
Cit3–
H+ 5b HCit2–
5a
4a
4b H2PO4–
Fig. 12.2 Physicochemical equilibria between milk components; Cit, citrate (adapted from Gaucheron, 2005).
is decreased (creaming of the fat owing to aggregation of the globules, variation in the salt composition, for instance), or owing to biochemical changes induced by active endogenous enzymes (proteases and lipases, for instance) and micro-organisms contained in milk, the best effect being the degradation of lactose and production of lactic acid causing a decrease in pH. 12.2.2 Whey Whey is the liquid co-product of cheese-making. At a first glance, it can be considered as milk without casein micelles and fat. Whey contains approximately 65 g L–1 dry matter: mainly lactose (~50 g L–1), nitrogen © Woodhead Publishing Limited, 2010
Separation technologies in dairy and egg processing 347 matter (proteins ~6 g L–1), ashes (minerals ~6 g L–1) and fat (0.3 g L–1) (Table 12.4). The composition of whey varies depending on the type of cheese produced and, more particularly, on the methods used for the coagulation of the milk. There are two main types of whey (Table 12.4): the ‘sweet’ whey is obtained by enzymatic hydrolysis of casein using rennet and its final pH is ~6.0–6.6; it contains the soluble glycomacropeptide portion of the k-casein that is released in the serum phase under the action of chymosin. The ‘acid’ whey is obtained after biological or chemical acidification of milk down to the isoelectric point of the caseins (pH 4.6) (Table 12.4). The acid whey has a higher mineral content than the sweet whey because of the release of minerals from the casein micelles (mainly calcium and phosphate) into the serum phase under acid conditions (Table 12.4, Fig. 12.2). Despite the complex composition of milk and whey, the main dairy constituents (fat, casein micelles, serum proteins, lactose and minerals) are well separated from others in terms of size (Fig. 12.1). The dairy fluids are therefore ideal starting materials for membrane separation processes although other industrial separation processes are developed on the basis of specific characteristics of milk constituents, such as differences in density, coagulation abilities and charge of constituents.
12.3 Pretreatment of milk using separation techniques 12.3.1 Skimming of milk and fractionation of fat globules Because manufacturers set a standard for their product and because most extraction procedures for milk proteins use skimmed milk as the starting fluid, Table 12.4 Approximate composition of sweet and acid wheys (from Marshall, 1982)
Sweet whey (g L–1)
Acid whey (g L–1)
Dry matter Nitrogen matter Non-nitrogen matter Lactose Ashes Fat Calcium Sulfate Magnesium Sodium Potassium Chloride lactate pH
66 6.2 0.37 52.3 5.0 0.2 0.5 0.7 0.07 0.53 1.45 1.02 – 6.4
64 5.8 0.40 44.3 7.5 0.3 1.6 0.5 0.10 0.51 1.40 0.90 6.4 4.6
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348 Separation, extraction and concentration processes standardization of the fat content of milk is classically applied in industrial manufacture. Separation of cream from skimmed milk has been a common practice for over 100 years, and is usually performed, at around 50 °C, by means of a flow-through centrifuge called a cream separator. As milk fat has a lower density than plasma, the fat globules rise under the influence of gravity, and the rate of rising is increased when a centrifugal field is applied and when the centrifuge is equipped with conical discs so as to limit the distance over which the fat globules have to move. Despite its widespread application in the dairy industry, this separation is not totally perfect and the skimmed milk, which generally represents 90% of the volume of the entering whole milk, still contains a very low fat content, 0.05 to 0.08%, which can strongly influence the efficiency of downstream processes and the resulting properties of dairy products. Moreover, this technology can not separate the native milk fat globules according to their sizes. In many dairy products, fat composition and structure cannot be adjusted easily because fat is usually homogenized. Homogenization reduces the fat globule size that results in small fat droplets that are disrupted and covered by caseins and serum proteins. Such modification of the fat globules characteristics promote their interactions with the cheese casein matrix which affect the structure of the rennet gel and decrease product functionality. Recently, separation of milk fat into small and large globules was proposed using ceramic microfiltration (MF) membranes under hydrodynamic conditions, which causes no damage to the native fat globule membrane (Goudédranche et al., 2000; Michalski et al., 2006). The process feasibility was shown but the development of industrial applications depends on the profits associated with the new products. It is claimed that the use of the small globule fraction in cheese production gives a smoother and finer texture, probably because of the interaction of fat globule membrane with the cheese casein matrix, and the differences in triglycerides content of the fat globules according to their size (Michalski et al., 2007). 12.3.2 Bacterial removal Centrifugation (CF) is also used to separate particles that have a density larger than that of milk serum, such as dirt particles, somatic cells, and micro-organisms. The bacterial quality of milk is the most variable factor with which the manufacturer has to contend. Pathogenic bacteria can contaminate milk and a wide variety of micro-organisms can lead to various cheese defects. Thus, in order to minimize the health hazard and control bacterial growth during milk processing, various combinations of time–temperature treatments (such as pasteurization 72 °C for 12 s) can be used, but they almost always affect the flavour, the functionalities of the milk components, and the cheesemaking properties. CF and crossflow MF are non-thermal preservation technologies that have been widely applied for removal of bacteria (Gésan-Guiziou, 2010).
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Separation technologies in dairy and egg processing 349 CF is mainly used for the removal of bacteria and spores from minimally pasteurized milk. It is especially applicable for spores removal, as their density is higher than that of bacteria, despite their smaller size. Its most important application has therefore proven to be for the removal of spores that undergo a late acid fermentation in the semi-hard and hard cheese process. It is also used to produce extended-shelf-life pasteurized milk, with a gain of about 3–5 days. Centrifugation is often referred as ‘bactofugation’ because the major manufacturer (Tetra Pak) produces centrifuges under the trademark of Bactofuge®. The separation is classically performed at 50–60 °C, which is a temperature similar to that in the cream separator, as the bactofuge is normally installed in series with the cream separator. At this temperature, the total bacteria count is reduced by 86.0–92.0% (decimal reduction of ~ 1 log), the spore removal reaches 90.0–98.0% (decimal reduction ranging from 1.0 up to 1.7), and a significant protein loss is observed, reaching 2.5 to 12% according to the types of machines. To reduce effectively the load of bacterial spores, a double bactofugation is practised in the cheese industry. An alternative to bactofugation is the use of MF, which was introduced successfully in the 1990s (Gésan-Guiziou, 2010). MF is particularly adapted to the removal of bacteria from skimmed milk, because the size ranges of fat globules and bacteria overlap (Fig. 12.1). The MF of whole milk would lead to a contaminated fat fraction which could be difficult to valorize. This separation is classically performed with a multichannel ceramic membrane with a pore diameter of 1.4 mm at a temperature of 35–55 °C: the milk components permeate through the membrane whereas the bacteria are satisfactorily retained. In order to overcome the membrane fouling phenomena, most MF plants in dairies operate using the hydraulic concept of uniform transmembrane pressure (UTP), that involves the circulation of the permeate co-current of the retentate (Alva-Laval Company, Bactocatch system) (Fig. 12.3) (Sandblom, 1974). In order to create a large pressure drop on the permeate side, this compartment is filled with plastic balls (Bactocatch system, Alva-Laval) or membranes are placed into small stainless-steel tubes in order to reduce the external space between the housing and the membrane porous media (Invensys APV). More recently, ceramic membranes with linear hydraulic resistance gradient were commercialized for this application: GP Membralox® membranes from Pall-Exekia (GP for permeability gradient); and Isoflux® membranes from Tami-Industries. This new membrane concept, with higher hydraulic resistance at the entrance of the membrane when the transmembrane pressure (TMP) is high, creates homogeneous filtration performance all along the membrane without a permeate circulation loop. By combining a high crossflow velocity (6–9 m s–1) and low TMP (~50 kPa), permeation fluxes are high (400–650 L h–1 m–2) for ~10 h, and matter losses are low: ~5% of the entering milk volume with a volume reduction factor (VRF), the ratio between either the volumes of feed and retentate in discontinuous mode, or the flow-rates of feed and retentate in continuous
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350 Separation, extraction and concentration processes
Permeate P
P
Retentate
Product
P
P
Fig. 12.3 Principles of the uniform transmembrane pressure system (Sandblom, 1974): the permeate circulates in co-current with the retentate in order to create a constant transmembrane pressure all along the filtering path. P, pressure sensor. Raw whole milk Skimming by centrifugation 50–60 °C
Cream
Skimmed milk 1.4 mm microfiltration 35–55 °C, VRF = 20
Retentate 1.4 mm microfiltration 35–55 °C, VRF = 10
Heat treatment
Heat-treated cream
Retentate
Microfiltered skimmed milk
Mixing, homogenization
Microfiltered whole milk
Fig. 12.4 Schematic representation of process for microfiltration of whole milk. Dashed lines, options for the treatment of retentate. VRF, volume reduction factor.
mode of 20, and ~0.5% with VRF = 200 (Fig. 12.4). The residual fat is largely retained (63%) and the proteins mostly transmitted (99%). The decimal reduction in bacteria count reaches 3–4 log with a high decimal reduction in spore count (2–4 log) that can be attributed to the binding of spores with cell walls. Somatic cells are also totally retained by the 1.4 mm
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Separation technologies in dairy and egg processing 351 MF membrane and, thus, the microfiltered milk is not degraded by their thermostable enzymes. MF is then much more efficient than bactofugation and makes it possible to decrease the microbial load of milk, while maintaining the organoleptic quality of milk owing to heat treatment at low temperatures (37–50 °C). The fat content of the microfiltered milk can be adjusted by addition of heat-treated cream (Fig. 12.4). The final retentate, that contains the bacteria and somatic cells in the milk, can be discharged separately for other suitable applications; blended with the cream and similarly heat-treated; or fed back to the cream for repeated separation. This process has been commercialized either for drinking or cheese milks. In most countries, mainly in Europe and North America (in particular Canada), the microfiltered milk intended for the drinking market undergoes a final pasteurization step to meet current regulatory requirements and lead to an extended life of 20–32 days at 4–6 °C. In cheesemaking, this process is used to produce safe raw milk cheeses, but requires knowledge about the microbial ecosystem composition that should be added to the treated milk for obtaining typical ripening of cheeses.
12.4 Standardization and concentration of milk proteins in the dairy industry The first milk protein concentrates were obtained by complexation of serum proteins with caseins after heat denaturation at ~90–95 °C (1–20 min), acid conditions (pH ~5.8–4.6) and possible addition of CaCl2. Precipitated products, referred to as ‘casein-whey protein co-precipitates’, were recovered with high yield (92–95%) reaching 97% with addition of calcium. However, after redispersion of the precipitates in the presence of a Ca chelatant (tripolyphosphate), proteins produced by this method had poor solubility properties. In the late 1960s, the development of membrane technology and more particularly of ultrafiltration revolutionized the dairy industry, offering the possibility of concentrating undenatured proteins. An ultrafiltration (UF) step was directly incorporated into the cheesemaking in what is known as the MMV process after the investigators Maubois, Mocquot and Vassal, who originally developed this process in 1969 (Maubois, 1981). Currently, UF is the most widely used membrane process for total protein concentration in cheese manufacture (Mistry and Maubois, 2004). The general idea behind this process is to concentrate all the milk proteins and simultaneously to remove the excess water and lactose by an initial UF step before coagulation, induced by addition of enzymes (rennet) and micro-organisms, thereby reducing or eliminating the need to separate the whey from the curd (Fig. 12.5). Spiral wound polymer membranes and to a lesser extent tubular ceramic membranes with a cut-off 10–50 kg mol–1 are used for this application, at a transmembrane pressure of 200–400 kPa and flux of 30–120 L h–1 m–2. © Woodhead Publishing Limited, 2010
352 Separation, extraction and concentration processes Traditional method
Ultrafiltration, MMV process
Milk
Milk Ultrafiltration
Rennet starters Coagulation
Pre-cheese Rennet starters
Coagulum
Draining
Cheese
Whey
Ultrafiltrate
Cheese
Fig. 12.5 Schematic representations of traditional ways of making cheese and the MMV process (with concentration to the final protein composition of the cheese).
Three categories of protein concentration level obtained by UF can be distinguished: standardization (retentates with VRF up to 1.7), intermediate concentration (1.7 < VRF < 5) and total concentration (VRF ~ 5–7). For almost 40 years, UF has been used in milk-processing plants, for total protein standardization (VRF up to 1.7) of both drinking and cheese milks. The main advantages of protein standardization, which results in a constant milk composition all year round independent of seasonal variations, are a more efficient use of the machinery and a better process control of the continuous cheesemaking processes. In addition, these concentrations provide economic benefits in terms of reduced requirement for coagulating enzymes and starters, as well as a slight increase in cheese yield, attributed to reduced losses of fat and casein particles in whey and better retention of serum proteins in the aqueous phase of cheese (Mistry and Maubois, 2004). An indirect but serious advantage of UF, admitted by the Codex Alimentarius, is the use of permeate to decrease the protein content of drinking milk to the minimum required value by law, 28 g L–1 (2.8% w/w) in most European countries. One can note that UF with VRF ~ 1.5 can also be incorporated into the manufacture of yoghurt and related fermented products, such as koumiss (Tamine and Robinson, 2007) in order to increase the protein levels without addition of milk powders. This technology is preferred to reverse osmosis because of the reduced level of lactose in the milk base. UF results in a better yoghurt gel strength compared with those obtained with addition of skimmed
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Separation technologies in dairy and egg processing 353 milk powders or concentration by vacuum evaporation. However, compared with cheesemaking, the increase of protein level is not accompanied in the yoghurt manufacture by a yield improvement, because the high-temperature heat treatment generally applied to the milk (typically 80 °C for 10 min) leads to the attachment of denaturated serum proteins with casein micelles, resulting in the concentration of serum proteins with casein micelles after heat treatment. Processing UF at higher concentration (VRF > 1.7) leads to a further advantage in terms of cheese yield. A distinction must be made between retentates concentrated to an intermediate level (1.7 < VRF < 5) in which some serum proteins are retained and syneresis still takes place, and concentration to the final composition of the cheese (VRF ~5-7), also called pre-cheeses, where the protein content is similar to the protein content of the final cheese and very little whey drainage is observed (Fig. 12.5). Intermediate concentrated retentates have been applied to numerous cheese varieties, ranging from soft to hard. Feta cheese manufacture, which is the main industrial application of this operation, yield increases of 14% while maintaining good quality products. The pre-cheeses concept was first applied to camembert cheese but many applications have been developed successfully for the manufacture of other cheese varieties, mainly fresh unripened cheeses such as quarg, ricotta and cream cheese, and soft and semi-hard cheeses including mozarella, Saint Paulin and feta, the manufacture of which being unquestionably the greatest success worldwide of the MMV process. The advantages of the MMV process are numerous. The overall cheese yield is about 10–30% higher than in the traditional process owing mainly to the retention of serum proteins and enzyme usage is generally reduced. The MMV process eliminates the need for the large storage tanks traditionally used for heating and cooking the curds, resulting in a saving in both capital investment and energy costs. In addition, it is possible to use the MMV process to convert much of the cheese production in a continuous operation, leading to significant advantages in terms of overall capital costs and operational efficiency. However, during UF, the concentration of serum proteins, with higher water-binding capacities than caseins, and of minerals associated with the casein micelles, that are lost in the whey during coagulation in a conventional process, led to the inherent properties of UF milk cheeses. Thus, rather than duplicating traditional cheese varieties, which requires numerous adaptations (Mistry and Maubois, 2004), UF technology leads to the development of new types of cheeses, well-accepted by consumers. For fifteen years, manufacturing technologies including UF have also emerged and have become important for the production of milk protein concentrates, presenting interesting new technical possibilities in cheesemaking. These concentrates contain both major milk protein groups in proportions similar to milk and are often dried. Because there are no specific standards
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354 Separation, extraction and concentration processes of identity for such concentrates, they cover a wide range of compositional and functional characteristics and are manufactured mainly using UF and diafiltration, for the reduction of the lactose content. Milk MF has also recently emerged for the production of protein concentrates with various casein–serum proteins ratio (12.5.2).
12.5 Isolation of whole casein in the dairy industry The characteristics of the two groups of milk proteins (casein and serum proteins) differ significantly, permitting their ready separation from each other by at least three techniques. 12.5.1 Isoelectric precipitation and rennet coagulation There are two principal ways of manufacturing casein on industrial scale: isoelectric precipitation and rennet coagulation, both based on the precipitation/ aggregation of casein. In isoelectric precipitation, destabilization of the casein micelles can be accomplished at pH around 4.6: when milk is acidified, the net charge of the micelles decreases, the calcium and phosphate are removed and the micelles become less and less stable until caseins precipitate. Acidification of the milk may be carried out by one of the following processes: – Inoculation of milk with a mixed or multiple defined strain starters, such as lactic acid-producing bacteria, which degrade some of the lactose to lactic acid during the period of incubation (about 14–18 h). – Direct addition of dilute mineral acid to the milk. Adding acid (HCl or H2SO4) to milk heated at 50 °C results in a casein curd (‘acid casein’) that can be separated by straining and washing several times. The coproduct, which is a mixture of acid whey and part of the washing water, has a high ash and chloride content (where HCl is used), which poses some difficulties during spray-drying. Acid destabilization of casein is exploited in the manufacture of cottage cheese, quarg and fermented milks. – Indirect acidification of the milk. A number of alternative processes have been proposed and patented to improve the quality of the acid whey. The acidification can be performed by contact of all or part of the milk with cation-exchange resins (strong hydrogen ion exchange resin), electrodialysis (Noël, 1992) or using carbon dioxide (Hofland et al., 2003). It is believed that these processes have relatively low commercial significance (Maubois and Ollivier, 1997). Regardless of the processes used, isoelectric caseins may be dried and used as such but they are insoluble in water, which limits their applications. Therefore caseins are usually converted to a soluble form, namely caseinate, © Woodhead Publishing Limited, 2010
Separation technologies in dairy and egg processing 355 by neutralization with alkali, usually NaOH (producing sodium caseinate) and less frequently KOH, NH4OH or CaOH2. Soluble caseinates may then be sprayed or roller dried. Destabilization of casein can also be caused by the proteolysis of the micelle stabilizing protein, namely k-casein. The chymosin, the principal enzyme present in calf rennet, releases the glycomacropeptide (C-terminal segments of glycosylated k-casein), and renders the casein micelles susceptible to precipitation with Ca2+ at the natural concentration in milk. This principle is used in the manufacture of rennet casein and most cheese varieties. 12.5.2 Microfiltration for fractionation of casein micelles and serum proteins The most promising technology for the selective separation of casein micelles is undoubtedly membrane MF using a membrane with a pore size of 0.1–0.2 mm. For the past 15 years, this operation has enjoyed rapid industrial development and is still expected to increase significantly over the next few years owing to the high quality and properties of the two fractions produced: – The retentate, enriched specifically in native casein micelles, improves the rennet coagulability in the cheesemaking process. Such casein enrichment reduces rennet coagulation time, accelerates curd firmness kinetics and increases final curd firmness compared with milk or caseinate. Consequently, casein and fat retentions into the cheese curd are significantly improved (less fines and fat in the drained whey) and that leads to an increase in cheese yield. The fact that a smaller amount of serum proteins end up in the cheese manufacturing process is also advantageous, because there are then fewer flavour and texture defects attributed to serum proteins and also fewer detrimental effects of heat treatment on rennet milk coagulability. The decrease in b-lactoglobulin (b-LG) content evidently lowers the extent of the formation of the complex k-casein/b-LG during heat treatment and, consequently, renneted micelles aggregate well. This property was used to develop a new high-heat milk powder, with a cheesemaking ability similar to that of raw milk (Fig. 12.6, Quiblier et al., 1991). – The permeate, with a composition close to that of a whey, contains soluble proteins in their native state, is sterile, and free of phage particles, cellular debris, fat and glycomacropeptide. It therefore becomes a useful fluid for preparing whey protein concentrates (WPC) and isolates (WPI) (protein ratio up to 97%) with very high functional and nutritional properties, in particular compared with WPC obtained from cheese whey (Foegeding et al., 2009). Despite these advantages, the microfiltrate is sometimes used in the cheesemaking process by reincorporation in the caseinenriched fraction after denaturation under moderate concentration and heat treatment (Fig. 12.6).
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356 Separation, extraction and concentration processes Skimmed milk MF 0.1 mm VRF = 3 Microfiltrate
Retentate Cream
UF 10–20 kg mol–1 VRF = 6
Cheese
Permeate
Concentration spray drying HT (95 °C, few minutes) Retentate WPC
Milk powder b-Lactoglobulin High cheese-making ability
Fig. 12.6 Primin process (Quiblier et al., 1991). VRF, volume reduction factor; WPC, whey protein concentrates; HT, heat treatment.
Finally, purified casein micelles obtained by diafiltration of retentate against water, and WPI obtained after UF of milk microfiltrate, are excellent starting fluids for further fractionation of individual caseins or serum proteins. From an engineering point of view, this separation is classically conducted using ceramic membrane either with the UTP system or membranes with linear hydraulic resistance gradient, as is done for bacterial removal (12.3.2). Separation is usually performed at 50–55 °C, at a crossflow velocity of 7 m s–1, TMP ~ 50 kPa and VRF of 2–4. At VRF = 3, serum proteins transmission is 65–80% depending on the milk heat treatment (pasteurization, thermalization or raw skimmed milk). Permeation flux is ª75 L h–1 m–2 for 10 h according to the critical stability criterion, which allows the industry to perform long runs with very moderate fouling and high selectivity (Gésan-Guiziou et al., 1999). Owing to the high running costs and investment required by the tubular ceramic equipment, the industry is starting to operate MF with spiral-wound polymer membranes. These membranes, which can not be used with the UTP system, are operated at low temperature (<10 °C) to avoid bacterial growth. Permeation flux is reported to be very low <10 L h–1 m–2 with a low protein transmission of 10–20%. This high serum protein rejection requires diafiltration against ultrafiltrate to decrease the soluble protein level in the enriched casein fraction. According to the manufacturers, about 25% of MF plants are equipped with organic membranes today and this proportion is likely to increase in the next years for cheese applications.
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Separation technologies in dairy and egg processing 357
12.6 Separation techniques applied to whey and derivatives in the production of cheese Because of the high water content and relatively high sugar and mineral contents, more than two thirds of whey production was disposed of as waste in the 1980s with a detrimental effect on the environment. Today, a significant amount of the whey produced is being processed, mainly for the recovery of its protein content. 12.6.1 Concentration of serum proteins It has long been the practice to obtain fully denaturated proteins by heating acidified whey (temperature >>90 °C and pH << 6). This method, derived from the technology used for making whey cheeses such as ricotta in Italy, results in a precipitate containing ~80% of the initial whey proteins. After centrifugation and spray drying, this protein preparation, called ‘lactalbumin’, was still impure and because of its high insolubility in water and very poor functionality found limited use (Pearce, 1992). The availability of UF membrane processes in the 1970s offered new possibilities to fully exploit the interesting nutritional, biological, functional and, in particular, solubility properties of the serum proteins. Whey has thus become to be regarded as a starting material for processing. Currently, two main processes have been industrially designed to produce whey protein concentrates, WPC containing 35–80% protein (expressed in nitrogen N ¥ 6.38 over dry matter), and whey protein isolates, WPI containing ≥ 90%. Most whey protein concentrates are produced using membrane technologies, and the most successful application of membrane processes in the dairy industry is the production of WPC by UF. UF membranes, with an appropriate cut-off (~10–20 kg mol–1) are used to remove both the lactose and ions, yielding a retentate with a high protein concentration, which can be further processed by evaporation and spray drying. The protein content of the final whey protein product depends on the degree of concentration during UF. For 35% WPC, a VRF of 4.5–7.0 is required, although it should reach 13–20 for 50–60% WPC. Combined with a diafiltration, which removes minerals and lactose from the retentate, whey UF (VRF 30–35) can lead to WPC purity of 75–85%. During UF, membrane fouling is mainly attributed to three different species: presence of residual lipids coming from the cheese manufacture, precipitation of calcium phosphate enhanced under neutral pH (7.0–7.5) and high temperature (55 °C) and accumulation of proteins at the membrane surface, more pronounced at pH close to their isolectric point (pH~ 5.0–5.5). Several whey pre-treatments, some using membrane operations, have been proposed to increase the purity of the final concentrates (particularly by reducing the residual lipids content) and to improve UF performance (in particular by limiting calcium phosphate precipitation and protein accumulation) (Maubois and Ollivier, 1997). Similar to WPC by
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358 Separation, extraction and concentration processes ultrafiltration, WPI can be produced using the microfiltration permeate of skimmed milk. In industrial processes, whey UF is performed mainly in multistage spiralwound systems with polyethersulfone membranes. Processes are currently operated at ª50 °C, requiring a pre-treatment in order to avoid severe fouling during operation. Flux is about twice as a high as flux at 10 °C, which is a major incentive for operating at such a high temperature. However, owing to the rapid decrease in membrane prices, a low temperature process (<15 °C) is now favoured to maintain the microbiological quality of the end product. Because serum proteins are amphoteric molecules, highly purified whey protein concentrates (WPI) are produced using ion-exchange chromatography, which provides an additional level of selectivity above membrane processing. To date, ion exchange is the main adsorption technique employed, but, with the advent of new technologies, there are potential applications of affinity binding and other techniques that have historically been economically limiting. At pH lower than their isoelectric point (pH ~5.0–5.5) proteins become positive and they can be adsorbed on cation exchangers, and at pH above they can be adsorbed on anion exchangers. After the adsorption of the proteins on the resins or columns (generally performed at pH ~3.2), lactose and other non-protein whey components are rinsed with water from the ion exchanger. Then alkali is added to a pH of about 8 to desorb proteins, and the desorbed proteins are eluted from the ion exchanger, concentrated by UF, evaporated and spray-dried. There are several ways of producing protein isolates, using, in particular, specific conditions of pH, ionic strength and ionic nature, which give the processors the opportunity to adjust the composition and functionality of the obtained fractions. Among them, two major ion exchangers are commercially available for this application: the Vistec and Spherosil processes. The Vistec process uses a cellulose-based cation exchanger, in a stirred tank reactor. WPI are then produced by a single-stage batch capture of proteins, whereas the Spherosil process uses a fixed-bed ion exchanger with porous silica beads coated with a polymer material having either cationic exchange potential (–SO3H groups, Spherosil S) or anionic exchange reactivity (–N(CH3)3 groups, Spherosil QMA). Acidified whey (pH <4.6) is generally applied to Spherosil S and sweet whey (pH 5.5) to Spherosil QMA. Using chromatography, protein concentrates are always devoid of lipids, which have high foaming properties. However, UF operation is, for economical reasons, the preferred method of processing. 12.6.2 Concentration and demineralization of whey and derivatives Concentration of whey and of various ultrafiltrates at their production site is the major application of reverse osmosis (RO) owing to its flexibility and energy consumption (9–20 kWh m–3 water removed) compared with vacuum evaporation (ª100 kWh m–3). © Woodhead Publishing Limited, 2010
Separation technologies in dairy and egg processing 359 RO, which is applied to remove water, is rarely applied to milk because the flux is small, the maximum concentration attainable is low, and this technology is not as attractive as UF in cheese and yoghurt manufacturing. However, RO of milk is sometimes used as a first step in the manufacture of milk powder to increase the capacity of evaporation plants and to reduce transport costs, as is commonly done for whey RO. RO is also used to concentrate ultrafiltrates, containing about 5% total solids, mainly lactose, salt and other minor soluble components of the milk. Concentrated ultrafiltrates have valuable uses such as animal feed, recovery of lactose after crystallization, fermentation of lactose to glucose and galactose as sweetener for confectionery industry, alcohol, and lactic acid. Concentration of whey by RO to a VRF of 4 and, hence, to about 25–28% dry matter is possible. Concentration is limited by high osmotic pressure, high retentate viscosity, calcium phosphate precipitation and lactose crystallization. The RO permeate can be reused for preparing cleaning solutions, but its composition is not similar to pure water. Urea can pass to some extent, and some salts and low-molar-mass peptides may do so. Because of the high salt content of whey, which generates numerous processing difficulties and nutritional imbalance (particularly in infant food), it becomes advantageous to demineralize whey before evaporation. The demineralization of whey can be achieved in various ways (electrodialysis, ion exchange, nanofiltration) according to the type of treated whey and the required demineralization level (Table 12.5). Regardless of the chosen processes, ions, and not undissociated salts, are removed. However, for a given overall proportion of salts removed, the rate of removal varies with the kind of ions and with the technology used. Nanofiltration (NF) for instance removes monovalent ions (such as Cl–) and concentrates divalent nutrition value ions such as calcium with the proteins. Whey is currently demineralized in the range 50–95% by electrodialysis and/or ion exchange, after being concentrated by RO, but these operations lead to large volumes of polluting effluents and high investment and running costs. NF is less efficient for salt removal than electrodialysis (which
Table 12.5 Demineralization of whey. Optimal processes as a function of whey type and demineralization level required
Demineralization level (% dry matter)
30%
50–70%
Sweet whey Nanofiltration Ion exchange + (6% dry matter) nanofiltration Concentrated Electrodialysis Electrodialysis whey (18–24% dry matter)
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>90% Electrodialysis + ion exchange + nanofiltration Ion exchange + electrodialysis
360 Separation, extraction and concentration processes can achieve 70% salt removal), but has the advantage of simultaneously concentrating the liquid, which is often desired, with low losses of lactose and nitrogen. Owing to the low osmotic pressure difference between retentate and permeate compared with RO (attributed to the transfer of monovalent ions), the TMP is lower and the operation is generally more cost-effective than electrodialysis. In addition, the NF step significantly improves the technological characteristics of the concentrate and gives it higher value (increase in yield of lactose crystallization, and low hygroscopy of obtained powders). In 10 years, nanofiltration has become the industrial method of choice for partial desalting of whey. It makes it possible to reach simultaneously the concentration of dry matter (20–22% at VRF ~4) and demineralization (25–50% and even 90% with diafiltration).
12.7 Fractionation of individual proteins and peptides in the dairy industry To a large extent, the properties of the WPCs and WPIs approximate the properties of b-LG, because it constitutes more than half of the whey proteins. To exploit the particular properties of individual proteins, which are known to exert a wide range of nutritional, functional and biological activities (Tables 12.2 and 12.6), and to emphasize the properties of b-LG, fractionation of Table 12.6 Milk protein functionality in foods (Maubois and Ollivier, 1997) Functional property
Mode of action
Food system
Solubility
Protein solvation
Beverages
Water adsorption and binding
Hydrogen bonding of water, entrapment of water
Meat sausages, cakes, bread
Viscosity Thickening, water binding
Soups, gravy, salad dressing
Gelation
Protein matrix formation and setting
Meats, curds, baked goods, cheese
Cohesion–adhesion
Protein acts as adhesive material
Meat sausages, baked goods, pasta products
Elasticity
Hydrophobic bonding in gluten, disulfide links in gels
Meats, bakery
Emulsification Formation and stabilization of fat emulsions
Sausages, salad dressing, coffee whitener, soup, cakes, infant formula
Fat absorption
Binding of free fat
Sausages, doughnuts
Foaming
Forms stable film to entrap gas
Chiffon desserts, cakes, whipped toppings
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Separation technologies in dairy and egg processing 361 whey protein mixtures for the isolation of one or a group of proteins is useful. Considerable progress has been made over the past twenty years in technologies aimed at separation, fractionation and isolation in a purified form of many interesting proteins occurring in both bovine colostrum and milk. Industrial-scale methods have been developed for some proteins but their large-scale manufacture is still limited. Chapter 16 gives more details on fractionation of individual proteins. 12.7.1 Fractionation of individual caseins There is a considerable interest in fractionating whole casein into individual caseins. Potential uses include bovine milk-based infant formulas and preparation of biologically active peptides and specific additives. The native casein micelles retentate obtained from skimmed milk 0.1 mm MF constitutes an excellent raw material for preparing individual caseins. Most studies on the fractionation of the whole casein are related to the isolation of b-casein, the main component of human casein, which contains numerous peptide sequences with high physiological properties. The isolation method is based on the preferential solubilization of the very hydrophobic b-casein at low temperature. At ~4 °C, b-casein dissociates from the casein micelle and can be isolated from the rest of the caseins (caseinate and renneted skimmed milk) using separation techniques such as UF and MF or centrifugation. The yield of b-casein is enhanced at low pH (4.2–4.6) and the b-casein purity can reach 90% (Le Magnen and Maugas, 1991). The co-product (retentate or sediment fractions) is enriched in as and k-caseins. Based on the same principle, a promising process to separate b-casein directly from whole milk has recently been proposed (Lucey and Smith, 2009). It operates in two successive MF steps both using polymeric spiral wound membranes. The first, operating at low temperature, separates b-casein from the retentate containing casein micelles and fat. b-casein is then separated from the rest of the materials present in permeate by warming the permeate to room temperature resulting in aggregation of b-casein. The second MF retain the aggregated b-casein while native soluble proteins passed into the membrane. 12.7.2 Fractionation of serum proteins Several methods related to the isolation of serum proteins have already been published, but many of them are only available at the laboratory scale (Bonnaillie and Tomasula, 2008). Some procedures have, however, led to industrial manufacture of enriched protein fractions or separate proteins such a-lactalbumin, b-lactoglobulin and minor proteins such as LF, LP and immunoglobulins. The cost of extracting these proteins is high but often justified when recognizing the great value-added benefits when incorporated in hygiene products, functional foods and nutraceutical products. Each of the extracted proteins or group of proteins has been proven or implied to have unique
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362 Separation, extraction and concentration processes functional, nutritional or nutraceutical properties. Some putative activities are digestive function (b-lactoglobulin), anticarcinogenic (a-lactalbumin), antimicrobial (LF and LP) and passive immunity (immunoglobulins). Among the commercially interesting proteins, the two main serum proteins, a-lactalbumin and b-lactoglobulin, can be produced in enriched fractions using membrane or centrifugation processes: a-lactalbumin has a great potential market because of its high content of tryptophan (four residues per mole) and in infant milk formula. The main uses of b-lactoglobulin appear to be in gel and foam-type products and in the manufacture of protein hydrolysates for food ingredients. Heat combined with pH adjustment can be exploited to fractionate b-lactoglobulin and a-lactalbumin from whey. a-Lactalbumin is a calcium metalloprotein (1 mol of calcium per mol of protein), that loses its bounded calcium and its stability when at ~55 °C (30 min) the pH is lowered to 3.8. At this pH calcium dissolves in the solution and a-lactalbumin unfolds and precipitates at 50–65 °C. Such physicochemical conditions involve the reversible polymerization of the protein that precipitates together with immunoglobulins and bovine serum albumin (Bramaud et al., 1997). Separation of the precipitate can be performed either by desludging clarifier or microfiltration. The supernatant, or permeate contains the b-lactoglobulin fraction that can be further processed by UF in combination with diafiltration to yield purified b-lactoglobulin (95% purity). The a-lactalbumin (60% purity) can be recovered from the sediment/retentate after solubilization at neutral pH, followed by UF. Starting from the permeate of milk microfiltration, this principle can be used to produce high purity non-lactosylated b-lactoglobulin (Maubois et al., 2001). Ion-exchange chromatography is also used industrially to produce a-lactalbumin and b-lactoglobulin enriched or purified (purity > 90%) fractions (Etzel, 1999; Etzel et al., 2006; Outinen et al., 1996). The high cost of purified fractions often prevents them from being used in targeted food applications. This process features the use of a resin to isolate fraction of protein from the rest of the whey. With a careful choice of the resin system and the eluants a-lactalbumin and b-LG can be separated from each other very precisely. For example, Thuran and Etzel (2004) modified a cationexchange chromatography method previously designed to fractionate sweet whey, to efficiently fractionate acid whey. They used different inexpensive food-grade buffers and the extraction was performed in two steps leading to the release of a-lactalbumin fraction with high recovery rate (96%) and purity (93%) on the one hand and WPI containing mainly b-lactoglobulin on the other hand. Recently, new technologies were shown to produce enriched soluble protein fractions (Bonnaillie and Tomasula, 2009). Supercritical CO 2 fractionation technology is one of the methods proposed for the recovery of large-scale serum protein fractions enriched with both a-lactalbumin and b-lactoglobulin from either WPC or WPI. When supercritical CO2 is injected © Woodhead Publishing Limited, 2010
Separation technologies in dairy and egg processing 363 into solutions containing soluble protein isolate or serum protein concentrate, fractions containing 70% w/w of a-lactalbumin in solid form and 95% w/w of b-lactoglobulin in a soluble liquid form (uncontaminated with chemical additives) were obtained. LF and LP are the two serum proteins having a high isoelectric point (Table 12.2) (Bonnaillie and Tomasula, 2008). Fat normally causes problems for chromatographic separations, as it blocks packed columns as soon as the feed is introduced. Therefore, fat is removed before cation-exchange capture of LF and LP from skimmed milk. They are then classically isolated by ionexchange chromatography from skimmed milk or whey. At the pH of milk or whey, both proteins are specifically adsorbed on cation exchangers, the other proteins being negatively charged. Their elution is generally realized through the use of an increasing ionic strength gradient. Recently, Andersson and Mattiasson (2006) extracted pure LF from WPC using a simulated moving bed chromatographic technology, with several advantages compared with non-moving bed columns (increase in productivity by 48%, increase in LF concentration by 6.5 times, reduction in buffer consumption by 4.3 times). The growing interest in LF is related to its antibacterial properties, by forming an iron complex and inhibiting the growth of micro-organisms by depriving bacteria of iron that is essential for their growth, and to its interesting nutritional activities owing to the transport of iron in the organism. Immunoglobulins can be isolated from whey using an UF membrane with a cut-off about 100 kDa or more but whey is a poor source of immunoglobulins compared with colostrum or milk produced by hyperimmunized cows (Table 12.2). Cow’s colostrum contains substantially higher concentrations of immunoglobulins than mature milk (20–200 g L–1 against 0.15–0.8 g L–1) and then can be used as a appropriate starting fluid for a two-step immunoglobulin extraction procedure (Piot et al., 2004): the colostrum is first microfiltered using a 0.1 mm pore membrane so as to obtain a permeate (named ‘serocolostrum’) which is crystal clear, free of blood and somatic cells as well as fat globules and casein micelles. The permeate that contains 80% of the initial immunoglobulins can then be further concentrated using ultrafiltration (100 kDa). Commercial immunoglobulin products are mostly used in veterinary medicine on neonatal ruminants and pigs. Because ruminants are born without blood antibodies, they are very susceptible to infection and it is highly desirable that they receive protection either by suckling colostrum for at least one week or by ingesting an immunoglobulin concentrate. 12.7.3 Fractionation of peptides Milk is a rich source of bioactive peptides and there is considerable commercial interest in producing bioactive peptides for use in food applications. Bioactive peptides are specific protein fragments, which have a positive impact on body functions and conditions and may ultimately influence health (regulation of weight; mood, memory and stress control; immune defence; and improvement
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364 Separation, extraction and concentration processes of heart, bone, dental and digestive health) (Korhonen, 2009) (Fig. 12.7). The market for bioactive peptides is increasing because the possibility of designing new dairy products with health-promoting benefits looks promising and offers a perspective for consumers and producers. Over the past few decades, a number of methods were developed for their purification (chapter 16) and, to-date, some casein-derived peptides have been manufactured on the industrial scale (Table 12.7). The most common way to produce bioactive peptides is through enzymatic hydrolysis of precursor proteins, using gastrointestinal enzymes, usually pepsin and trypsin. After hydrolysis, the peptides are fractionated and enriched using various methods (precipitation with salts or solvents, ultrafiltration or chromatography). Angiotensin-converting enzyme (ACE)
Egg white pH ª 8.5
Extraction of lysozyme (cation exchange chromatography)
Egg white free of lysozyme
Three steps: • fixation • rinsing (water) • elution (NaCl)
Extracted lysosyme
Crystallization of lysozyme (pH ª 10)
Concentration (filter press)
Solubilization (pH ª 3.5–5.0)
Clarification, concentration, drying (centrifugation, ultrafiltration, filter press, spray drying)
Lysozyme powder
Fig. 12.7 Lysozyme extraction using ion exchange chromatography (adapted from Guérin-Dubiard and Anton, 2010).
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Separation technologies in dairy and egg processing 365 Table 12.7 Marketed bioactive peptides (from Korhonen, 2009) Brand name
Type of product
Claimed functional bioactive peptides
Health/function Manufacturer claims
Calpis Sour milk Val-Pro-Pro, Ile-Pro- Reduction Pro, derived from of blood b-casein and k-casein pressure Evolus Calcium- Val-Pro-Pro, Reduction of enriched Ile-Pro-Pro, derived blood pressure fermented from b-casein milk drink and k-casein Biozate Hydrolysed b-Lactoglobulin Reduction whey protein fragments of blood isolate pressure BioPURE- Whey protein 106–109 fragment Prevention of GMP isolate of k-casein dental caries, influence the clotting of blood, protection against viruses and bacteria PRODIET Flavoured aS1-casein f(91-100) Reduction of F200/ milk drink, (Tyr-Leu-Gly-Tyr-Leu- stress effect lactium confectionery, Glu-Glu-Leu capsules Leu-Arg) Festivo Fermented aS1-casein f(1–6), No health low-fat (1–7), (1–9) claim hard cheese Cysteine Ingredient/ Milk protein- Aids to raise peptide hydrolysate derived peptide energy level and sleep C12 Ingredient/ Casein-derived peptide Reduction of hydrolysate blood pressure Capolac Ingredient Caseinophosphopeptide Helps mineral absorption PeptoPro Ingredient/ Casein-derived peptide Improves hydrolysate athletic performance and muscle recovery Vivinal Ingredient/ Whey protein derived Aids relaxation alpha hydrolysate and sleep Recaldent Chewing Calcium casein Anticariogenic gum peptone–calcium phosphate
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Calpis Co., Japan Valio Oy, Finland
Davisco, USA
Davisco, USA
Ingredia, France
MTT Agrifood Research, Finland DMV International, the Netherlands DMV International, the Netherlands Arla Foods Ingredients, Sweden DSM Food specialist, the Netherlands
Borculo Domo Ingredients (BDI), the Netherlands Cadbury Adams, USA
366 Separation, extraction and concentration processes inhibitory peptides, known for their antihypertensive property, and calciumbinding phosphopeptides, for example, are commonly produced by trypsin (Korhonen and Pihlanto, 2006). For the preparation of phosphopeptides from casein, Brulé et al. (1981) proposed the use of an ultrafiltration membrane for processing permeate after digestion of caseinate in solution with a proteolytic enzyme. The separation of the phosphopeptides present in the permeate was performed by UF of the peptide solution after addition of a bivalent cation salt (calcium chloride) so as to cause aggregation of phosphopeptides. The non-phosphorylated peptides pass through the membrane and diafiltration against water, used to purify the phosphopeptides in the retentate, results in a preparation which is rich (>90% w/w) in the desired phosphopeptides. The glycomacropeptide, the C-terminal part of the k-casein released in whey by the action of chymosin was shown to be separated from sodium caseinate using UF membrane or centrifugation. This peptide has numerous uses, such as action on satiety and inhibition of Escherichia coli cells adhesion to intestinal walls and, in particular, it contains no phenylalanine, which makes it suitable for use as a nutritional protein supplement for patients suffering from phenylketonuria, who cannot digest protein containing phenylalanine owing to their lack of the appropriate degrading enzyme. A number of dairy starter cultures, as well as proteolytic enzymes isolated from lactic acid bacteria also lead to the formation of bioactive peptides from milk proteins, in particular during the manufacture of fermented dairy products. For example, several studies have demonstrated that Lactobacillus helveticus strains are capable of releasing antihypertensive peptides. A UHT milk fermented by the GG strain of Lb. rhamnosus and subsequently digested with pepsin and trypsin produced hydrolysate fractions that were immunosuppressive. The occurrence of many bioactive peptides in bovine milk is now well established (Korhonen, 2009), but currently the industrial-scale production of such peptides is limited by a lack of suitable separation technologies. Among them, membrane techniques, such as NF or UF, are used industrially to produce ingredients that contain bioactive peptides based on casein or whey protein hydrolysates (Table 12.7) and seem to be the best technology available for the enrichment of bioactive peptides.
12.8 Treatment of effluents and technical fluids in the dairy industry In many countries, the dairy industry is considered to be one of the largest generators of food-processing wastes. Owing to large water volumes for factory cleaning and disinfection, the dairy industry produces 0.2–11.0 L effluents/L of processed milk. Cleaning-in-place (CIP) operations significantly contribute to water consumption and are responsible for 50 to 95% of the overall volume of waste streams (Daufin et al., 2000, Alvarez et al., 2007). © Woodhead Publishing Limited, 2010
Separation technologies in dairy and egg processing 367 The polluting load of these effluents ranges from 0.2 to 5.0 g L–1 of chemical oxygen demand (COD), and is mainly caused by loss of raw material (0.5 to 2% of milk). CIP systems are also mainly responsible for the high pH value (9–11) of the end-of-pipe wastewaters. Although a high portion of dairy wastewaters is still land spread, the treatment of the effluents mainly takes place in biological treatment plants at most dairies. For a factory of average capacity (106 L of milk a day), the sludge produced is usually used for land spreading (1 to 3 tonne of dry matter) and purified water drained to rivers (0.3 ¥ 106 to 3 ¥ 106 L). Current European regulations relating to landfill management, land spreading and purified water quality along with social pressure, has recently forced the dairy industry to significantly reduce its production of sludge and to improve purified water quality. Over the past few years, the dairy industry has been attempting to find new concentration and separation processes to reduce its effluents production. Several types of effluents are currently treated in the industry, benefiting mainly from the potential of membrane technology: washing waters of rennet casein precipitate were treated, respectively, using NF and dead-end filtration; white (flushing) and pre-rinsing waters, corresponding to the first step of CIP, were treated using UF, NF or RO (Blanchard, 1991; Delbecke, 1981). The outcome is a highly significant improvement of water quality after treatment [suspended solids (SS) < 2 mg L–1; chemical oxygen demand (COD) <35 mg L–1; BOD5 <3 mg L–1) and the re-use of milk components (either as recycling back to the production unit or animal feed). Evaporation condensates (also called ‘cow’s water’) and permeates of milk and whey NF with COD of 10–1000 mg L–1, can be treated by RO with eventual UF or MF pretreatment (Horton, 1997). In both cases, the RO permeate produced can be used as a source of water with ‘food quality’, for rinsing and cleaning operations. Cheese brines (170–230 g kg–1 NaCl) are widely recycled after a UF (50 kDa) but more commonly after a MF (0.2–1.4 mm) step, which strongly reduce microbiological counts, without altering the chemical composition in contrast to conventional pasteurization (Pedersen, 1992). Heat treatment and Kieselguhr (diatomaceous earth) filtration are still the most well-used technologies for brine treatment but the former modifies the calcium phosphate equilibrium in solution, and the latter is recognized by The World Health Organization as a cause of lung disease, thus requiring safe working conditions. Finally, used alkaline and acid CIP solutions, which are periodically drained to waste, can be advantageously regenerated. The automatic renewal of cleaning solutions once a week, which is generally practised, leads to an excessive consumption of water and chemicals, and mainly caustic soda (about 120 tonne/year for a plant producing 106 L day–1 of milk). Desludging of caustic soda solutions is currently practised, but MF, which retains suspended solids and transmits small molecules responsible for the low surface tension of the re-used solutions, seems to be an appropriate operation both from the
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368 Separation, extraction and concentration processes technical and the economical points of view (Gésan-Guiziou et al., 2007). From the beginning of the cleaning process and regardless of the equipment used and the type of CIP processing, low values of surface tension (attributed to chemical reactions of proteins and fat in alkaline conditions) were observed in caustic soda solutions when their loading charge was increased (Alvarez et al., 2007). Because the low surface tension characteristics led to high cleaning kinetics (Alvarez et al., 2007), the MF-regenerated caustic soda solutions is as fast as that of a commercial alkaline detergent, and much more efficient than fresh caustic soda.
12.9 Conclusions and future trends in the dairy industry Membrane separation technologies and, to a lesser extent, chromatography, offer to the dairy technologist several techniques for the extraction and purification of almost all the main proteins of milk and for the treatment of wastewaters. Membrane operations in particular have made it possible for new and original products to be created. The fractionation of casein micelles from serum proteins using MF membrane is one of the best examples of such a success, which would have been considered unlikely some years ago. The differences in attributes between serum proteins obtained from milk (using milk MF) and those isolated from cheese whey are becoming an important factor for dairy protein processors considering, very recently, this new avenue for fractionating proteins and for producing bioactive peptides. Apart from being a balanced source of valuable amino acids, milk proteins contribute to the specific properties of various dairy products, and possess interesting biological properties, which make them potential ingredients of health-promoting foods. A few commercial protein and peptide fractions have been launched on the market and this trend is likely to continue alongside increasing knowledge about the functionalities of the products. Advances in membrane design, the best understanding of the limiting phenomena occurring during filtration operations, and the recent decrease of the processing cost of polymer membrane, make this technology more and more attractive. Therefore further integration of membrane operations is to be expected, provided they are designed in such a way that at each processing step, membrane fouling is limited, membrane cleaning is optimized, and end products, co-products and wastes are given equal attention. Through the dairy industry, membrane processes have been shown to provide the food industry with efficient tools for limiting the environmental impact of the food sector. There is no doubt that in the near future any food process will include at least one membrane operation for their effluents treatments. New applications are also likely to be developed, such as recovery of phospholids derived from the fat-globule membranes from buttermilk (aqueous phase produced from churning butter) and recovery of growth factors present in cows’ colostrum. Moreover, new emerging technologies, © Woodhead Publishing Limited, 2010
Separation technologies in dairy and egg processing 369 such as rotating and vibrating filtration, emulsification, and supercritical carbon dioxide fractionation, have met with some success in the laboratory and this could lead to more commercial applications in the future.
12.10 The egg products industry and composition of egg products Owing to their high functionalities (foaming, coagulative, emulsification, and binding), many egg products are used as ingredients in various food applications, such as bakery products, meringues, mayonnaise, cookies and meat products. Egg white in particular is in great demand in the baking, confectionery and cake mix industries for its whipping and foaming properties and it is mainly used in the form of dried egg white solids. Moreover, eggs are an excellent source of high-quality proteins containing all nine essential amino-acids and other nutrients required in our diets, and possessing bioactive properties including antimicrobial activity, protease inhibitory action, immunomodulatory, anticancer and antihypertensive activities, vitamin binding properties and antigenic or immunologic characteristics. Knowledge of the characteristics of the individual proteins has led to the development of new separation processes now industrially used for the extraction of three proteins of egg white, and of some components of the yolk. Eggs are composed of three main parts (Burley and Vadehra, 1989): the eggshell with the eggshell membrane, the albumen also named egg white, and the yolk. The yolk is surrounded by the white, which in turn is enveloped by eggshell membranes and finally by a hard eggshell. The egg white makes up about 66% of the liquid weight of the egg. It contains about 88–90% water. Proteins are the major components of albumen solids (about 10–11% of the white weight) whereas carbohydrates (mostly free glucose) (≈ 0.8–1.0%), minerals (≈0.5%) and lipids (≈0.03%) are minor components. Table 12.8 lists the content and some characteristics reported for some proteins in egg white. It should be noticed that despite the numerous recent studies for separating and identifying the proteins located in hen’s egg, many proteins remain uncharacterized or even unknown. Except for lysozyme and avidin, most proteins are acid proteins and are negatively charged at the natural pH of the egg white (pH 9.0–9.3). Glucose is the main ‘free’ sugar, and is usually removed by bacterial fermentation and enzyme hydrolysis before drying of egg white. In some applications, the presence of glucose is undesirable because it causes a detrimental effect on storage stability and quality of the product, by the forming of off-flavours and brown pigments caused by Maillard reaction. Egg yolk contains about 48–51% water, according to the age of the laying hen and the duration of preservation. Lipids are the main components (31% of the total weight) of the egg yolk solids. The lipid distribution is 65% © Woodhead Publishing Limited, 2010
370 Separation, extraction and concentration processes Table 12.8 Content and characteristics (isoelectric point, molecular weight) of some proteins found in egg white (from Li-Chan and Kim, 2008) Protein
Protein % of Isoelectric point egg white
Ovalbumin 54 Ovalbumin Y Ovotransferrin 12 Ovomucoid 11 Ovomucin 3.5 Lysozyme 3.4 Ovoglobulin G2 globulin 4.0 G3 globulin 4.0 Ovoinhibitor 1.5 Ovoglycoprotein 1.0 Ovoflavoprotein 0.8 Ovomacroglobulin 0.5 Cystatin 0.05 Avidin 0.05
4.5 (5.1–5.3) (5.3–5.5) 6.1 (6.2–6.7) 4.1 4.5–5.0 10.7 (6.1–5.3) 5.5 4.8 5.1 (6.2–6.4) 3.9 (5.0–5.4) 4 (5.0–5.2) 4.5 5.1 (6.1) 10
Molecular weight (kg mol–1 or kDa) 45 (42.4) (53.4–54.3) 76 (85–75) 28 (37.2–43.1) 5500–8300 14.3 (15) 30–45 Not determined 49 (69.5–63.6) 24.4 (37.2–43.1) 32 (37.4–40) 769 12.7 (17) 68.3
Source: data were compiled from Li-Chan et al. (1995), except for those in parentheses, which are from Guérin-Dubiard et al. (2006).
triglycerides, 28–30% phospholipids, and 4–5% cholesterol. Yolk contains about 16% proteins, consisting of livetins (globular proteins) and lipoproteins particles including low and high-density lipoproteins. As a food, yolk is a major source of vitamins and minerals (3.5% of dry yolk). Whole egg contains about 25% solids, 23% proteins and 10% fat. Minor amounts of minerals and carbohydrate are also present.
12.11 Concentration and stabilization of egg white and whole egg The concentration of egg white and whole egg (it is not necessary to concentrate egg yolk, which contains about 50% solids) is used in two main industrial applications: the concentration of egg components before spray drying and the stabilization of egg white and whole egg. Owing to its high water content, the egg white is quasi-systematically concentrated before spray drying to reduce energy costs. Egg products are rarely concentrated using conventional thermal evaporation because of the damage caused to the highly heat- and shear-sensitive egg white proteins. Membrane processes, especially RO and UF with low molecular weight cutoff (MWCO) in order to limit the loss of nitrogen matter, were investigated as a means of concentrating egg before drying (Bergquist, 1995). These processes have been commercially applied to the concentration of egg white.
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Separation technologies in dairy and egg processing 371 RO concentrated all egg components and led to a permeate, free of organic matter, that could be re-used for intermediate cleaning of operations. UF is sometimes more useful because it not only removed water, but also lowered glucose, sodium and potassium levels by about 50% in the concentrated product. The partial removal of glucose reduced problems observed during storage and excessive browning during baking processes. A higher removal of glucose, while retaining the egg white protein, can be achieved by carrying out UF in conjunction with diafiltration. UF led also to higher permeation fluxes, lower energy consumption (owing to lower applied TMP), superior functional properties of the concentrates owing to the removal of free glucose and salts. Neither of the concentration methods (RO or UF) significantly affected foaming properties, and UF was observed to improve gel strength of egg white, which was probably related to the increase in protein concentration. The concentration of whole egg using UF is marginal. The concentration of liquid egg white and whole egg is used for their stabilization. The objective of this operation is to concentrate the egg product by UF and then add salt and/or sugar to decrease the water activity (aw) of the product under the thresholds of micro-organisms development (aw = 0.85) (Bonduelle, 1978; Liot 1980). This process enables the processing of concentrated egg white up to 33% dry matter content with a maximum sugar content of 50% or salt content of 9% (0.80
12.12 Industrial extraction of egg-white proteins Owing to the high nutritional, functional and biological properties of eggwhite proteins, many studies are currently in progress for the extraction of
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372 Separation, extraction and concentration processes albumen proteins. Several methods of proteins isolation have already been proposed, but very few have been developed on an industrial scale (Nau et al., 2010). Lysozyme, ovotransferrin and avidin are the three proteins commercially removed from egg white. 12.12.1 Lysozyme Lysozyme is the only hen’s egg protein routinely used commercially. It can be used as a preservative and finds practical applications in the food, pharmaceutical and medicine sectors (Lesnierowski and Kijowski, 2007). For example this protein is used in cheese to prevent contamination because it does not inhibit the starter and secondary cultures required for the ripening of the cheeses. It demonstrates antibacterial properties, particularly against Gram-negative bacteria, among them a number of food pathogens such as Listeria monocytogenes. Lysozyme is known as a hydrolysate that cuts the b-1-4 linkage of the glycosidic bond between polysaccharide copolymers, which represent structural units of many bacterial cell walls. Lysozyme from hen’s egg white is a small protein with a molecular weight of 14.2 kDa having, apart from the other egg albumen proteins, a strong basic character (isoelectric point 10–11, Table 12.8). Both physical and chemical properties of lysozyme have been exploited as isolation processes. The small size of the protein was used in membrane techniques, particularly UF or MF, to separate lysozyme from egg white. However, the ability of this enzyme to electrostatically bind with other negatively charged egg white proteins greatly reduced the transmission of the protein through the membrane. Despite these membrane processes being tested on a large scale (Peri and Feriscini, 1972, Lepienne et al., 1986, Kijowki et al., 1998), they are not yet developed on an industrial scale. Currently, the commercial procedures of lysozyme isolation focus on two main processes, based on the very high isoelectric point of lysozyme: the selective precipitation of lysozyme and extraction using ion-exchange chromatography. The selective precipitation of lysozyme combines adjustment of the pH to 10 (leading to a pH close to the isoelectric point of the protein) and addition of 5% sodium chloride to increase ionic strength and favour precipitation. With this technique, several precipitations and resolubilizations are necessary to obtain high-purity protein. The remaining egg white after lysozyme separation becomes difficult to valorize owing to the high residual content of salt. It is possible to reduce its salt content by using UF and diafiltration techniques and desalted remaining egg white has been shown to maintain foaming properties when compared with the native egg white. Ion-exchange chromatography and, in particular, cation-exchange chromatography has mainly been used on an industrial scale to isolate lysozyme. In association with salt precipitation of the extracted lysozyme, this technique can lead to a very pure protein fraction (>99%) and to a co-
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Separation technologies in dairy and egg processing 373 product with high techno-functional properties. Figure 12.7 gives the main steps of the industrial process: after slight acidification of the egg white down to 8.5, the egg white was put into contact with ion-exchange resins, such as Amberlite, carboxymethylcellulose (CMC), carboxymethyl-Sephadex, or Duolite. Both column and batch techniques are used for large-scale operations. In the batch system, resin is stirred so as to maintain it in suspension and favour the sorption of lysozyme from egg white. When the lysozyme is fixed on the resin, the egg white free of lysozyme is extracted; the resin is washed with water before the elution of the protein with sodium chloride solution. Crystallization of lysozyme, initiated with already formed crystals, is then performed at a pH close to 10 in order to increase its final purity. The final lysozyme crystals are then concentrated by press filter, solubilized at acid pH (3.5–5.0), clarified by centrifuge, concentrated by UF, filtered again with a press filter for bacterial removal and finally dried. 12.12.2 Ovotransferrin Ovotransferrin is produced on an industrial scale, but to a lesser extent than lysozyme (Guérin-Dubiard and Anton, 2010). This protein is similar to serum transferrin in animals having the functions and interesting nutritional activities as a result of the transport of iron in organisms. Many procedures, mainly based on liquid chromatography as for all proteins belonging to the transferrin family, have been developed to purify this protein. Owing to its neutral isoelectric point (pI 6.2–6.7, Table 12.8) the extraction can be performed either by cation or by anion-exchange chromatography. When performing cation-exchange chromatography, lysozyme should be first extracted. When performing anion-exchange chromatography, all the proteins with an isolectric point lower than the ovotransferrin one should be first separated. Yields of extraction vary from 50 to 80%, and ovotransferrin purity can reach 98% (Guérin-Dubiard and Anton, 2010). One can note that the extraction of ovotransferrin can be performed from egg-white free of ovomucin which can easily be obtained after centrifugation of the egg-white after dilution with water (1:4) and adjustment of the pH at 6. Ovomucin has the ability to precipitate at low ionic strength, which leads to serious problems during water washing phases of the chromatography process. The extraction of ovotransferrin from the egg-white free of ovomucin can lead to a protein purity up to 90%. 12.12.3 Avidin Avidin represents a maximum of 0.05% of the total protein content of eggwhite (Table 12.8). It is best known for its binding properties with biotin (vitamin H or B8), an essential growth factor. Owing to its high biotin affinity, avidin is thought to serve as a defensive protein against biotinrequiring micro-organisms. It is also widely used as tool in a number of
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374 Separation, extraction and concentration processes affinity-based separations, in biochemical diagnostic assays, and in a variety of other applications. Several isolation processes based on chromatographic separations have been proposed either based on ion-exchange or affinity properties. Avidin is an alkaline glycoprotein similar to lysozyme (Table 12.8), and then its isolation can be performed using processes based on ion- (mainly cation-) exchange chromatography. However, the extracted avidin is often contaminated by lysozyme. Several authors proposed adapted procedures in order to increase the selective elution of avidin fixed on the cationic support (Nau et al., 2007). Durance and co-workers proposed a novel ion-exchange chromatography procedure depending on pH, ionic strength and/or nature of the salts used in the eluting buffer (Durance and Nakai, 1988; Durance et al., 1991). This procedure made it possible to simultaneously extract avidin and lysozyme from undiluted egg white. Avidin recovery was around 75% and the purity rate of the avidin could reach 40%. The purity of avidin was increased by a subsequent chromatography step. Rao et al. (2003) proposed another original procedure: the selective elution of avidin fixed on a Streamline SP (sulfopropyl) support was performed using hydroxyazobenzene-2’-carboxylic acid (HABA). Avidin, which possesses a high affinity for this molecule, is the only protein being eluted, leading to a very high purity (98%). Owing to its high biotin affinity, avidin is also purified using affinity chromatography. The high commercial value of this protein makes it acceptable to extract this protein with the costly but very selective affinity processes. The recovery and purity of the protein obtained are very high (> 95%), but the materials are very expensive and their lifetimes are rather limited, which could explain why affinity chromatography is classically not used for large-scale separations. Garret-Flaudy and Freitag (2000/2001) proposed an affinity precipitation process using iminobiotin–polymer, which gave an avidin purity higher than 90%.
12.13 Industrial extraction of yolk components The fractionation of yolk components does not exclusively concern the isolation of proteins, as previously seen for the egg white. The g-livetin (egg antibody which corresponds to the immunoglubulin of yolk, immunoglubin Y, IgY), and the phospholipids mainly composed of phosphatidylcholine, PC (80–85% commonly named ‘lecithin’), and phosphatidylethanolamine, PE (10–15%), are commercially extracted. g-Livetin has been exploited on an industrial scale for nearly twenty years because its extraction is easier than immunoglobulins from mammalian blood. This egg antibody is promising for use in immunoassays, to quantify toxins or pathogenic viruses for example, and its use as a functional tool in pharmaceutical applications is developing. Various extraction methods were reviewed in detail by Schade et al. (2005). Two main isolation procedures © Woodhead Publishing Limited, 2010
Separation technologies in dairy and egg processing 375 have been developed. The first method comprises the precipitation of the g-livetin by polyethyleneglycol and the second one uses ion-exchange chromatography leading to 70% recovery with 60% purity. In both cases, the protein is first recovered in the soluble phase of the yolk by water dilution (6 times; pH 5, 4 °C for 6 h). Egg-yolk phospholipids offer several industrial applications, mainly in the nutritional, pharmaceutical and cosmetic fields. This fraction is ultimately used in the food industry as an emulsifier, viscosity reducer, lubricant, and as an antispattering, wetting and release agent, and its widespread uses are linked to its ability to act as a surface-active substance in multiphase systems. Phospholipids from yolk could also serve as encapsulation systems (liposomes and double emulsions) intended for cosmetic formulas and medical uses. They have several other properties including antioxidative activity and inhibition of cholesterol absorption. Industrial extraction methods for phospholipids are commonly based on organic solvents. Phospholipids are soluble in hydrocarbons and other organic solvents, whereas they are typically insoluble in acetone. The latter characteristic allows the separation of the accompanying lipids in order to increase the purity of the phospholipids. Ethanol is also used to increase PC and PE because they are highly soluble in it. On this basis, Juneja et al. (1994) proposed a method for large-scale preparation of phospholipids, treating fresh egg yolk by a combination of acetone and ethanol additions and filter press separations. This process, which produced phopholipid fraction containing the same proportion of individual phospholipids as in the initial yolk (80–85% PC; 10–15% PE), can be extrapolated to the industrial scale (Guérin-Dubiart and Anton, 2010; Juneja et al., 1994). However, because the use of organic solvents in common extraction procedures is questionable on the grounds of safety, alternative techniques for example using supercritical carbon dioxide fluid have been studied and gave positive results (Sim, 1994).
12.14 Conclusions and future trends in the egg-processing industry The egg-processing industry has made a great deal of progress since the 1950s, but current research by food scientists is still adding to the understanding of the components of eggs and their fractionation. The recent isolations on a laboratory scale and characterizations of egg components give clear evidence of the potential uses of eggs: the protein and lipid fractions of egg white and yolk not only provide interesting functional properties but also offer real opportunities for nutraceutical and human health applications. Efforts are still required to identify and produce molecules of interest for exploiting the potential of egg-derived molecules.
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376 Separation, extraction and concentration processes The egg-processing industry has recently learned to fractionate their products. Currently lysozyme is the only egg component routinely removed from egg. The separations used to produce ovotransferrin, avidin, IgY and phospholipids enriched fractions are likely to be developed, but more potential utilization of egg can be envisaged, mainly in biotechnology processes (chapter 19). Enzymatic hydrolysis of egg white to produce various peptides is promising. Peptides have bacteriostatic or antioxidant activities; some lower blood pressure of hypertensive rats. Thus the egg-processing industry has a strong base to expand into new innovative applications, with particular emphasis on the heath benefits of eggs. Finally, for environmental reasons, the treatment of the wastewater produced by the egg-processing industry is likely to develop in the next few years. The effluents issued from the eggbreaking machines for instance is particularly high with a BOD5 ranging from 1–22 g L–1. The membrane technologies will certainly offer good opportunities to decrease the load of effluents sent to the purification plant, in a similar way to that previously done in the dairy industry.
12.15 Sources of further information and advice Two clear and comprehensive books, which emphasize basic aspects of all kinds of membrane processes, but do not give much detail on dairy and egg products applications: ∑ ∑
Zeman LJ and Zydney AL (1996), Microfiltration and ultrafiltration: principles and applications, New York, Marcel Dekker; Cheryan M (1998), Ultrafiltration and microfiltration handbook, Lancaster, Technomic Publishing.
More detailed information about principles of cheesemaking and general aspects of dairy science in: ∑
Walstra P, Wouters JTM and Geurts TJ (2006), Dairy science and technology, Boca Raton, FL, CRC Press, Taylor and Francis.
More detailed information of membrane processes applied to dairy fluids in: ∑
Britz TJ and Robinson RK (2008), Advanced dairy science and technology, Oxford, Blackwell Publishing; ∑ Mistry VV and Maubois JL (2004), ‘Application of membrane separation technology to cheese production’, in Fox PF, McSweeney PLH, Cogan TM and Guinee TP, Cheese chemistry, physics and microbiology. Vol 1 General aspects, London, Elsevier.
More detailed information about egg composition and functionalities in: ∑
Huopalahti R, Lopez-Fandino R, Anton M and Schade R (2007), Bioactive egg compounds, Berlin, Springer; © Woodhead Publishing Limited, 2010
Separation technologies in dairy and egg processing 377 ∑
Mine Y (2008), Egg bioscience and biotechnology, Hoboken, John Wiley & Sons Inc.
More detailed information about egg processing in: ∑
Nau F, Guérin-Dubiard C, Baron F and Thapon JL (2010), Science et technologie de l’œuf et des ovoproduits. Vol 1: Production et qualité de l’œuf; Vol 2: De l’œuf aux ovoproduits, Paris, Tec&Doc Lavoisier (in press).
12.16 References Ahmad S, Gaucher I, Rousseau F, Beaucher E, Piot M, Grongnet JF and Gaucheron F (2008), ‘Effects of acidification on physico-chemical characteristics of buffalo milk: a comparison with cow’s milk’, Food Chem, 106, 11–17. Alvarez N, Gésan-Guiziou G, Daufin G (2007), ‘The role of surface tension of reused NaOH on the cleaning efficiency in dairy plants’, Int Dairy J, 17, 404–411. Andersson J, Mattiasson B (2006), ‘Simulated moving bed technology with a simplified approach for protein purification – separation of lactoperoxidase and lactoferrin from whey protein concentrate. J Chromatogr A, 1107(1–2), 88–95. Bergquist DH (1995), ‘Egg dehydration’, in Stadelman WJ and Cotterill OJ, Egg science and technology, Haworth Press, Fourth edition, Binghamton, NY, 335–369. Blanchard BD (1991), ‘Plant effluents dairy waste streams recovery’, Dairy Food Environ Sanit, 11(9), 494–496. Bonnaillie L, Tomasula PM (2008), ‘Whey protein fractionation’, in Onwulata CI and Huth PJ, Whey processing, functionality and health benefits, Wiley Blackwell, Singapore, 15–39. Bonnaillie L, Tomasula PM (2009), ‘Supercritical carbon dioxide fractionation of serum protein isolate for new food-grade ingredients’, IFT ‘09 Annual Meeting & Food Expo. Paper No. 251: 14. Bonduelle M (1978), ‘Une nouvelle technique dans la conservation des ovoproduits’, Ind Agric Aliment, 95, 1043–1048. Bramaud C, Aimar P, Daufin G (1997), ‘Whey protein fractionation: isoelectric precipitation of a-lactalbumin under gentle heat treatment’. Biotechnol Bioeng, 56(4), 391–397. Brulé G, Roger L, Fauquant J, Piot M (1981), ‘Phosphopeptides from casein-based material’, US patent, 4358465. Burley RW, Vadehra DV (1989), The avian egg: Chemistry and biology, New-York Wiley Interscience. Daufin G, Gésan-Guiziou G, Boyaval E, Buffière P, Lafforgue C, Fonade C (2000), ‘Minimization des rejets liquides de l’industrie laitière par traitement des effluents à l’aide de procédés à membrane’, Tribune de l’eau, 600(4), 25–33. Delbecke R (1981), ‘Recovery of milk by hyperfiltration’, Milchwissenschaft, 36(11), 669–672. Durance TD, Nakai S (1988), ‘Simultaneous isolation of avidin produced by binding of biotin’, J Food Sci, 53, 1096–1106. Durance T, Li-Chan E, Nakai S (1991), ‘Process for the isolation and separation of lysozyme and avidin from egg white’, Canadian Patent 1,283,072. Etzel MR (1999), ‘Isolating b-lactoglobulin and a-lactalbumin by eluting from cation exchanger without sodium chloride’, US patent 5,986,063. Etzel MR, Helm TR, Vyas HK (2006), ‘Methods and compositions involving whey protein isolates’, US patent, 60,569,078.
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378 Separation, extraction and concentration processes Foegeding EA, Zulweska DM, Barbano DM, Drake MA, Luck PJ, Yong YH, Vardhanabhuti B, Berry T (2009), ‘Comparison of the functional properties of serum proteins isolated from milk or whey’ J Dairy Sci, 92(E-Suppl. 1), 163. Froning GW (2008), ‘Egg products industry and future perspectives’, in Mine Y Egg bioscience and biotechnology, Wiley Interscience, Hoboken, New Jersey, US. Garret-Flaudy F, Freitag R (2000/2001), ‘Use of avidin-(imino) biotin system as a general approach to affinity precipitation’, Biotechnol Bioeng, 71, 223–234. Gaucheron F (2005), ‘The minerals of milk’, Reprod Nutr Dev, 45, 473–483. Gésan-Guiziou G (2010), ‘Removal of bacteria, spores and somatic cells by centrifugation and microfiltration techniques’, in Griffiths M Improving the safety and quality of milk, Woodhead Publishing Ltd, Cambridge. Gésan-Guiziou G, Alvarez N, Jacob D, Daufin G (2007), ‘Cleaning-in-place coupled with membrane regeneration for re-using caustic soda solutions’ Sep Purif Technol, 54, 329–339. Gésan-Guiziou G, Boyaval E, Daufin G (1999), ‘Critical stability conditions in crossflow microfiltration of skimmed milk: transition to irreversible deposition’, J Membr Sci, 158, 211–222. Goudédranche H, Fauquant J, Maubois JL (2000), ‘Fractionation of globular milk fat by membrane microfiltration’ Lait, 80, 93–98. Guérin-Dubiard C, Anton M (2010), ‘Fractionnement de l’œuf’, in Nau F, Guérin-Dubiard C, Baron F and Thapon JL Science et technologie de l’œuf et des ovoproduits Vol 2: De l’œuf aux ovoproduits, Lavoisier Tec&Doc, Paris. Guérin-Dubiard C, Pasco M, Molle D, Desert C, Croguennec T, Nau F (2006), ‘Proteomic analysis of hen egg white’. J Agric Food Chem, 54(11), 3901–3910. Hofland GW, Berkhoff M, Witkamp GJ, Van der Wielen (2003), ‘Dynamics of precipitation of casein with carbon dioxide’. Int Dairy J, 13(8), 685–697. Horne DS (2006), ‘Casein micelle structure: models and muddles’. Curr Opinion Colloid Interface Sci, 11, 148–153. Horton BS (1997), ‘Water, chemical and brine recycle or reuse – applying membrane processes’ Austr J Dairy Technol, 52(1), 68–70. Juneja LR, Sugino H, Fujiki M, Kim M, Yamamoto T (1994), ‘Preparation of pure phospholipids from egg yolk’, in Sim JS and Nakai E, Egg uses and processing technologies – new developments, CAB International, Wallingford, UK, 139–149. Kijowki J, Lesnierowski G, Fabisk-Kijowska A (1998), ‘Methods of lysozyme separation, enzyme molecular form and functional quality of the residual egg white’, in The 2nd international symp egg nutrition newly emerging ovotechnologies, Banff, Alberta, 54. Korhonen H (2009), ‘Milk-derived bioactive peptides: from science to applications’ J Funct Foods, I, 177–187. Korhonen H, Pihlanto A (2006), ‘Bioactive peptides: production and functionality’ Int Dairy J, 16, 945–960. Korhonen H, Pihlanto A (2007), ‘Technological options for the production of healthpromoting proteins and peptides derived from milk and colostrum’, Curr Pharm Des, 13, 829–843. Le Magnen C, Maugas JJ (EURIAL) (1991), ‘Method and device for obtaining beta casein’, Patent PCT / FR 91/00506. Lepienne A, Maubois JL, Thireau M, Piot M (1986), ‘Procédé pour l’obtention de lysozyme par microfiltration à partir d’une matière à base de blanc d’œuf’, French patent 25697222. Lesnierowski G, Kijowski J (2007), ‘Lysozyme’, in Huopalahti R, Lopez-Fandino R, Anton M and Schade R Bioactive egg compounds, Springer, Berlin, pp. 33–42. Li-Chan ECY, Kim HO (2008), ‘Structure and chemical composition of eggs’, in Mine Y Egg bioscience and biotechnology, Wiley Interscience, Hoboken, New Jersey, USA 1–95.
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Separation technologies in dairy and egg processing 379 Li-Chan ECY, Powrie WD, Nakai S (1995), ‘The chemistry of eggs and egg products’, in StadelmanWJ and Cotteril OJ, Egg science and technology, Haworth Press, Fourth edition, Binghamton, NY, 105–175. Liot R (1980), ‘Produit hautement concentré de blanc d’œuf ou d’œuf entier salé et son procédé de préparation’, French patent 8022309. Lucey J, Smith K (2009), ‘An integrated processing system to produce beta-casein, native serum protein and casein concentrates from whole milk’, J Dairy Sci, 92(ESuppl.1), 164. Marshall KR (1982), ‘Industrial fractionation of milk proteins: serum proteins’, in Fox PF Developments in dairy chemistry-1, Applied Science Publishers, NY. Maubois JL (1981), ‘Perspectives d’utilisation des techniques à membranes dans les industries agro-alimentaires’. Académie d’Agriculture de France, Procès-verbal 26 nov. 1980, 1451–1461. Maubois JL, Fauquant J, Famelart MH, Caussin, F (2001), ‘Milk microfiltrate, a convenient starting material for fractionation of whey proteins and derivates’, in Proceeding of the 3rd International Whey Conference, Munich Sept 12–14, Behr’s Verlag, Hamburg. Maubois JL, Ollivier G (1997), ’Extraction of milk proteins’, in Damodaran S and Paraf A, Foods proteins and their applications, New York, Marcel Dekker Inc, 579–595. Michalski MC, Camier B, Gassi JY, Briard-Bion V, Leconte N, Famelart MH, Lopez C (2007), ‘Functionality of smaller vs control native milk fat globules in Emmental cheeses manufactured with adapted technologies’, Food Res Int, 40, 191–202. Michalski MC, Leconte N, Briard-Bion V, Fauquant J, Maubois JL, Goudédranche H (2006), ‘Microfiltration of raw whole milk to select fractions with different fat globule size distributions: process optimization and analysis’, J Dairy Sci, 89(10), 3778–3790. Mistry VV, Maubois JL (2004), ‘Application of membrane separation technology to cheese production’, in Fox PF, McSweeney PLH, Cogan TM and Guinee TP, Cheese chemistry, physics and microbiology. Vol 1 General aspects’ London, Elsevier. Mulder H, Walstra P (1974), ‘The milk fat globule. Emulsion science as applied to milk products and comparable foods’, Commonwealth Agricultural Bureaux, Farnham Royal, UK. Nau F, Guérin-Dubiard C, Baron F, Thapon JL (2010), ‘ Science et technologie de l’œuf et des ovoproduits’ Vol 2: De l’œuf aux ovoproduits, Paris, Tec&Doc Lavoisier (in press). Nau F, Guérin-Dubiard C, Croguennec T (2007), ‘ Avidin’, in Huopalahti R, Lopez-Fandino R, Anton M and Schade R Bioactive egg compounds, Springer, Berlin, 75–80. Noël R (Société Vidaubanaise d’Ingénierie) (1992), ‘Procédé de séparation du phosphocaséinate de calcium et du lactosérum d’un lait écrémé plus généralement d’un composé protéique d’un liquide biologique’. PCT WO 92/12642. Outinen M, Tossavainen O, Syväoja EL (1996), ‘Chromatographic fractionation of a-lactalbumin and b-lactoglobulin with polystyrenic strongly basic anion exchange resins’, Lebens Wis Technol, 29(4), 340–343. Pearce RJ (1992), ‘Protein recovery and whey protein fractionation’. in Zadow JG Whey and lactose processing, Elsevier Applied Science, 271–316. Pedersen PJ (1992), ‘Microfiltration for the reduction of bacteria in milk and brine’ In: ‘New applications of membrane processes’, Int Dairy Fed Bulletin, Special issue 9201, 33–50. Peri C, Feriscini C (1972), ‘Concentration and fractionation of egg white by ultrafiltration’ Sci Technol Alimenti, 2, 120–122. Piot M, Fauquant J, Madec MN, Maubois JL (2004), ‘Preparation of “serocolostrum” by membrane microfiltration’, Lait, 84, 333–342. Qi P X (2007), ‘Studies of casein micelle structure: the past and the present’, Lait, 87, 363–383. Quiblier JP, Ferron-Baumy C, Garric G, Maubois JL (1991), ‘Procédé de traitement des
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380 Separation, extraction and concentration processes laits permettant au moins de conserver leur aptitude fromagère’, Patent FR 2 681 218 A1. Rao M, Gupta M, Roy I (2003), ‘Process for the isolation and purification of a glycoprotein avidin.’ International patent 03/099035 A1. Sandblom RM (Alfa-Laval) (1974), ‘Filtering process’, Swedish patent 7,416,257. Schade M, Calzado EG, Sarmiento R, Chacana PA, Porankiewicz-Asplund J, Terzolo HR (2005), ‘Chicken egg yolk antibodies (IgY-technology): a review of progress in production and use in research and human and veterinary medicine’. ATLA, 33, 129–154. Sim JS (1994), ‘New extraction and fractionation method for lecithin and neutral oil from egg yolk’, in Sim JS, and Nakai E, Egg uses and processing technologies – new developments CAB International, Wallingford (UK), 128–138. Stadelman WJ, Cotteril OJ (1995), Egg science and technology, 4th ed. Binghamton, NY: Haworth Press. Tamine AY, Robinson RK (2007), ‘Yoghurt: science and technology’. CRC Press, Abington, England. Thuran KN, Etzel MR (2004), ‘Whey protein isolate and a-lactalbumin recovery from lactic acid whey using cation-exchange chromatography’, J Food Sci, 69(2), 66–70.
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Separation technologies in the processing of fruit juices 381
13 Separation technologies in the processing of fruit juices G. Vatai, Corvinus University of Budapest, Hungary
Abstract: After an introduction to the fluids in the fruit juice product sector, the various separation processes for production of fruit juices from different fruits are discussed. The separation processes for the production of fruit juice concentrate are reviewed, and the advantages and disadvantages of the applied techniques are discussed. Finally, a multistep membrane process for must concentration is described using laboratory experimental data, modeling and optimization. Key words: fruit juice, separation, extraction, membrane filtration, membrane distillation, osmotic distillation.
13.1 Introduction With the changing of nutrition trends, interest has focused on plants and crops with many valuable components. Fruits have always played a very important role in this. Fruits can be consumed fresh in their natural form, even in winter (e.g. apples and oranges) when they are stored in the proper way. Before the 20th century, drinking squeezed fruit juices was the privilege of rich people. Nowadays, in order for us to consume some kinds of fruits all year round, it is necessary to produce concentrate with a high level of dissolved solids for storing the concentrate frozen, in the refrigerator or at room temperature, depending on the nature of the dissolved solid content. Concentrate can be diluted by water, or used to create a fruit juice with the same characteristics as the original fresh one, if the process of concentration has been carried out correctly. The industrial concentration of fruit juices is usually performed by
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382 Separation, extraction and concentration processes multistage vacuum evaporators. In most cases, the volatile components are recovered and added back into the concentrated product at a later stage. Evaporation often causes heat degradation of many valuable compounds. To prevent this significant decrease in quality, it is necessary to employ nonthermal methods of concentration, such as freeze concentration systems and membrane processes. The most common membrane techniques are the membrane filtration processes. Micro- (MF) and ultrafiltration (UF) are used as clarifying processes, whereas nanofiltration (NF) and reverse osmosis (RO) are used for pre-concentration of the juices (Lagana et al. 2000, Nene et al. 2002, Rektor et al. 2006, 2007).
13.2 Characteristics of foods/fluids in the fruit juice product sector In the world market there are countless fruit juice based products. They differ mainly in terms of raw material, composition, fruit content, nutritional value, sensory characteristics and packaging, but in some cases the biggest difference is the brand name (Hui et al., 2006). Fruit juice based drinks are classified on the basis of fruit content, into three categories: ∑ juices and fruit musts; ∑ fruit nectars; and ∑ soft drinks with fruit content. Juices and fruit musts are produced by mechanical procedures, mainly pressing, and the produced juice has the same taste, color and aroma as the original fruit. The final fruit juice composition is also identical to the original fruit. Juices cannot contain food additives (preservatives, aromas and coloring agents). Juices are consumed in their fresh form soon after production, or used as a raw material for concentration. Fruit juices can be divided into two subcategories: they can be filtered, i.e. clarified to be transparent (apple, grape), or they are cloudy, containing colloids and fibers like all citrus-based juices. Fruit nectars are made from fruit pulps or fruit juices diluted with sugar syrup. They usually contain only one fruit, such as apple, orange or peach, but they can be made from blends of more than one fruit juice or pulps. The preparation of blends and minimum fruit content are regulated by government standards, industrial specifications and other requirements (Hui et al., 2006). In order to ensure international trade, these standards conform to the Codex Alimentaria of the FAO/WHO Food Standard Program. The raw materials used in the production of fruit juices can be grouped as follows: citrus, pomaceous fruits, stone fruits, grape and berries. Some of the raw materials, such as, apple and orange are suitable for juice production
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Separation technologies in the processing of fruit juices 383 the produced juice can be consumed without any additive, but the juice of some berries (sea buckthorn, black and/or red currant) are very acidic and are enjoyable only when blended with sugar syrup or concentrate of apple juice or grape must.
13.3 Designing separation processes to optimize product quality in the fruit juice product sector Juice extraction, the removal of juice from fibrous solid particles, is the basic operation of fruit juice production. The fruit has to be prepared for this operation, i.e. the separation step. This preparation step depends on the type of fruit. In some instances, the fruit has to be chopped (apple, peach), whereas in other instances (cherry, sour cherry, plum, apricot) it is very important to remove the stem before chopping. A typical flow sheet of the natural juice and juice concentrate production is shown in Fig. 13.1. In this scheme of concentrate production, several conventional separation processes can be recognized: ∑ mechanical pressing of the juice from the fruit pulp; ∑ juice extraction from the marc by water as solvent; ∑ clarification of the fruit juice by centrifugation or filtration; ∑ concentration of the fruit juice by multistep vacuum evaporation. The fruit juice production technology presented in Fig. 13.1 is a typical one (Barta and Körmendy, 2007) for the production of fresh natural fruit juice and concentrate for longer storage. In the past few decades, some of the conventional separation processes have been replaced by newer ones, for example the clarification step, where the traditional method of clarification by centrifugal separation or filtration has been replaced by ultrafiltration, especially in the case of clarified or ‘transparent’ juice production, such as apple and grape (Barta and Körmendy, 2007). In this traditional technology, the juice yield can be improved by using enzyme treatment for pectin degradation, as well as through the extraction of the residual fruit juice from the marc using hot water (50–90 °C) (Hui et al., 2006). 13.3.1 Juice extraction by pressing For juice extraction by pressing, the liquid content of the fruits is separated from the solid particles. The most common method of this separation is a mechanical pressing of the juice out from the fruit pulp. The type of equipment utilized in this separation depends on fruit species. Where the hard parts of the fruit (e.g. the stem) have been previously removed (cherry, plum, apricot) it is possible to use typical mechanical pressing with higher pressure, whereas for pressing berry-type fruits, such as grapes and currants, mild pneumatic
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384 Separation, extraction and concentration processes Reception of the fruits
Cleaning of the washing water
Washing
Stem elimination Marc Juice extraction–pressing Secondary juice extraction
Aroma extraction
Clarification (centrifugation)
Enzyme treatment
Pasteurization
Clarification (centrifugation)
Storing
Concentration (evaporation)
Natural, unfiltered juice
Storing
Fruit juice concentrate
Fig. 13.1 Production of natural fruit juice and fruit juice concentrate using conventional separation processes.
pressing is more effective (Barta and Körmendy, 2007). In both instances, the pressing processs requires outside forces to create tension in the system and drain out the liquid, which results in some shape modification. In batch systems, the solid phase will stay in the pressing vessel, whereas the liquid will be drained out across a sieve and/or filter media. The remaining solid, with low liquid content, is called marc. The most important parameter of the pressing process is the liquid or ‘juice yield’, which refers to the percentage of juice pressed out, compared with the amount of raw material that was entered into the system. The juice yield is determined basically by the preparation and pretreatment (enzymatic hydrolysis or not) of the fruit before pressing, and the pressure applied (Barta and Körmendy, 2007; Hui et al., 2006). In the basket type batch pressing machines commonly used in the fruit juice processing industry, press machine volume is often decreased by mechanical or hydraulic forces; the space available for the fruit pulp/marc may also be
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Separation technologies in the processing of fruit juices 385 decreased by pneumatic interaction at lower pressures, up to a few bars, to press the juice out of the fruit. 13.3.2 Juice extraction using water as solvent Fruit juices can also be produced by solid–liquid extraction. This process can be characterized by the degree of extraction, expressing the amount of valuable substances extracted, compared with the total valuable matter content of the fruit. In this operation, mass transfer is controlled by the molecular diffusion of certain components, which can be improved by increasing the temperature of the operation. In order to increase the diffusion coefficient and the permeability of the cell walls, this solid–liquid extraction is performed at 50–70 °C. The mass transfer can be improved by increasing the specific area available for transport, in most instances by decreasing the particle size of the solid phase or by chopping the fruits. Another means of improving the mass transfer is increasing the concentration gradient for the mass transfer using a higher solvent–solid ratio, multistep cocurrent or countercurrent operation mode (Fonyó and Fábry, 1998). Diffusion juice extraction is usually carried out in double-screw extractor devices (Hui et al., 2006). 13.3.3 Juice clarification The extracted fruit juices are usually turbid, containing plant particles (fibers, cellulose, starch, and lipids) and colloids such as pectin, proteins, and polyphenols. Depending on the nature of the final product, these substances must be partially or totally removed to avoid further turbidity and precipitation and to improve sensory attributes such as taste, color, and odor. This clarification step can be performed by physicochemical or mechanical methods, as well as by using combinations of these methods. The physicochemical methods of clarification are those in which clarifying agents and enzymes are added during the procedure. For the clarification of fruit juices, mineral clarifying agents (bentonite and silicic acid), natural organic (gelatine) or organic polymers (polyvinypolypyrrolidone) are often used. The similarity between these clarifying agents is that all of them are charged (Barta and Körmendy, 2007), a characteristic which causes the bentonite to adsorb proteins, and the gelatine and polyvinylpolypyrrolidone to precipitate negatively charged particles (polyphenols and decomposed pectin). During juice clarification, the pectin molecules have to be decomposed, because they hinder aggregate formation and the settling of floating substances. Besides pectin hydrolysis, the decomposition of starch and proteins can be carried out on the same time scale using pectin, which results in the application of a mixture of enzymes (Hui et al., 2006). The aim of mechanical clarification is also to remove suspended solids using only mechanical separation processes like centrifugation or filtration. This separation process is performed in settling centrifuges, eliminating fibers
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386 Separation, extraction and concentration processes from cloudy juices. Filtration is the next step of fruit juice production. The traditional method is using devices based on a slurry layer and an operation mode using silica or perlite as additives. The filtration equipment should be frame filter press, continuous vacuum drum filter or others. Ultrafiltration and microfiltration have been widely applied in the clarification of filtered juice membrane filtration in the past two decades (Fonyó and Fábry, 1998; Bélafi Bakó, 2002). These membrane filtrations are able to solve the problem of clarification and filtration in one step. These clarified and filtered or cloudy juices are then ready for consumption. The juice can be packed and pasteurized and released to market, or it can be used as a suitable raw material for the production of semifinished products.
13.4 Production of fruit juice concentrate As shown in Fig. 13.1, after production (either by pressing or solid liquid extraction) and certain post treatments, the fruit juice can be concentrated for longer shelf life, decreasing the water content and increasing the total soluble solid (TSS) content. This concentration improves storage and transportation properties and costs, but it has to be carried out very carefully in order to avoid the loss of aroma components and valuable ingredients, and to minimize changes in the sensory properties of the juice. The most common method of juice concentration is evaporation (Barta and Körmendy, 2007). 13.4.1 Concentration by evaporation The most common method of fruit juice concentration is evaporation, and, from a physical point of view, this means the water evaporating from the boiling liquid phase. This process is carried out in devices such as an evaporator and steam provides the energy for boiling and water evaporation. However, valuable fruit components are mostly heat sensitive, so, for fruit juice concentrate production, vacuum evaporation is recommended (Ashurst, 2005). To decrease energy consumption during this process, batteries of 3–4 evaporator elements are commonly used (Fonyó and Fábry, 1998). The chemical, rheological and thermal characteristics of juices play an important role in this evaporation–condensation process. Because these parameters tend to vary from juice to juice, the operation parameters are different for different fruits or juices. The type of evaporators chosen for an individual juice is based on the characteristics of that particular juice. The most widely used type of evaporators are the film, pipe, plate, and centrifugal based ones (Barta and Körmendy, 2007). Evaporators are usually combined with aroma recovery units, in most cases a distillation column. The condensed aroma components are often remixed into the concentrate to improve the flavor,
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Separation technologies in the processing of fruit juices 387 or can be concentrated and applied as natural aroma extracts in other fruit products. 13.4.2 Concentration by freezing During freezing, the fruit juice is cooled down under water at 0 °C at atmospheric pressure. The water forms ice crystals and these ice crystals are separated from the suspension, thus lowering the water content, i.e. the concentration of TSS will be higher. Owing to the high cost of producing ‘cold energy’, this technology is only appropriate for the concentration of valuable and heat-sensitive fruit juices. This concentration method can be achieved in one or several steps, and, with this technology, the eutectic concentration of the fruit juice–water system can be achieved (Barta and Körmendy, 2007). The separation of the solid phase (ice) and the liquid phase (concentrated juice) should be carried out by mechanical separation: settling, filtration, and centrifugation. Only clarified juices should be concentrated using this technology, because for unclarified juice separation, the fibers and colloids are removed together with the ice. On the other hand, this type of concentration is a very gentle process, because there are no aroma, color or vitamin losses during this operation, owing to the low temperatures involved. Its disadvantages are high energy consumption and lower concentration efficiency compared with evaporative concentration (Barta and Körmendy, 2007). 13.4.3 Concentration by membrane separation processes The sterilization of fruit juice can be achieved without heating by using membrane separation processes, by the mechanical removal of microbes during clarification. On the other hand, water removal (concentration of fruit juice) can be achieved at or near room temperature (Jiao et al., 2004). For the clarification of the juice, ultrafiltration (UF) or microfiltration (MF) is suitable (De Carvalho et al., 2008; Cassano et al., 2006, 2007; He et al., 2007). Even when we use ultrafiltration we are removing more valuable components (polyphenols) than is ideal. For the concentration of fruit juices, reverse osmosis (RO) (Rektor et al., 2004; Kozak, 2005) and/or nanofiltration (NF) can be used (Kiss et al., 2004; Porter, 1990). The only problem is that during this concentration process on the retentate side of the membrane, the concentration of the TSS becomes higher and higher, causing higher osmotic pressure and decreasing the driving force of the concentration process, which is the difference between the transmembrane pressure (TMP) and osmotic pressure of the retentate (fruit juice concentrate) and permeate (practically water). Using the usual TMP (40–50 bar), the fruit juice can be concentrated up to 23–26% TSS (Belafi-Bako et al., 2000; Belafi-Bako, 2002). The advantages of this concentration technology include better quality juice concentrate as valuable components like aroma, vitamins, polyphenols are concentrated and
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388 Separation, extraction and concentration processes not destroyed, and, on the other hand lower energy consuption and lower cost in comparison with evaporation, as well as with freeze concentration (Ashurst, 2005; Kiss et al., 2004). Having used membrane technology to reach the final concentration (45–65% TSS) for storing the concentrate, at room temperature or little bit lower, the semi-product produced by reverse osmosis (23–26% TSS) should be further concentrated by membrane distillation (MD) (Laeson, 1997; Lagana et al., 2000; El-Bourawi et al., 2006) osmotic distillation (OD) (Lefebre, 1988; Courel et al., 2000; Thanedgunbaworn et al., 2007) or nanofiltration (NF) (Kiss et al., 2004; Porter, 1990; Vatai, 2007). From the previous statements, the fruit juice concentration by membrane technology could be solved in a three-step process: clarification (UF or MF), the production of ‘half-concentrate’ by a pressure-driven membrane process (RO), and the production of ‘final concentrate’ using three different methods (Vatai, 2007). One of the most valuable fruits in Hungary is the grape. As an example, the production technology for the must concentration is shown in Fig. 13.2, supported by experimental and optimization data (Vatai, 2007). The conditions for grape growing in the country are very good, and a large volume and good quality must is produced. Most of the vintage is processed by the viticulture and soft drinks industry. Because the ripening of the grape and the vintage happen just once a year, there can sometimes be a lack of grapes (or poor quality) or an overproduction. The preservation of grape juice can solve these problems. Viticulture can produce must concentrate from the surplus that can then be applied to the process of upgrading poorquality, low-sugar content grape juice, or the soft drinks industry can develop a new product for its customers. This technology for the production of must concentrate has advantages in terms of saving both energy and quality. However, it also has some other benefits in terms of food safety and the impact on the environment, because this is practically a clean technology:
Colloids, microorganisms
Must 12–16% TSS
MF
~Water
Clarified must 12–16% TSS
Concentrate 23–26% TSS
RO
NF
MD
OD
Concentrate 45–60% TSS
Concentrate 55–65% TSS
Concentrate 55–65% TSS
Fig. 13.2 Alternative methods of grape juice (must) concentration by complex membrane processes.
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Separation technologies in the processing of fruit juices 389 the main product is the must concentrate, the byproducts are the membrane filtered water and the concentrate of the clarification, which can be added to grape marc for ‘pálinka’ or spirit production, but can also be used as raw materials for the production of valuable products rich in antioxidant capacity by conventional or supercritical extraction. In the proposed must concentration technology, the first step is clarification by microfiltration. The aim of this step is: ∑
the removal of suspended solids, which causes a decrease in the viscosity and results in a higher filtration capacity, ∑ in further membrane filtration processes for must concentration, the clogging of the circulation channels, especially in spiral wound membrane filtration elements, is minimized, ∑ the micro-organisms are also removed, resulting in the sterilization of the must. Figure 13.3 contains typical experimental data of changing permeate flux during the clarification of typical Hungarian musts Kékfrankos and Furmint, by microfiltration. From Fig. 13.3 it can be seen that, after a certain period of time, the flux decreasing rate has been stabilized, reaching a steady-state flux. From the diagram it is also obvious that there are some differences between the permeate fluxes of white grape must (Furmint) and blue grape must (Kékfrankos), owing to the greater total solid content of the Kékfrankos must (Vatai, 2007). From the clarified must, the so-called ‘half concentrate’ can be produced by reverse osmosis, with TSS between 23 and 26%. This can be stored frozen for a long time, thus maintaining the valuable component concentration. The results of the concentration of Furmint must in a pilot scale RO plant in the laboratory of Food Engineering, Corvinus University of Budapest, 70 Furmint
Flux, J (L m–2 h–1)
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1.6
1.8
Fig. 13.3 Permeate flux changing during clarification of typical Hungarian musts, Furmint and Kékfrankos, by microfiltration.
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390 Separation, extraction and concentration processes is presented in Fig. 13.4. The flux of the permeate decreases during this process because of the lower driving force caused by the decreasing osmotic pressure of the concentrate, as well as by the increasing viscosity. From the diagram, it is also obvious that the final concentration at room temperature (50 bar transmembrane pressure and 600 L h–1 volumetric recirculation flow rate) was slightly more than 23% TSS and, at the end of the experiment, the permeate flux decreased almost to zero (Vatai, 2007). The ‘half-concentrate’ produced can be stored frozen for even longer but the storing costs, as well as the transportation costs of the semi-product from the producer to the drink producer, would be quite high if the concentration was produced near to the location where the fruit was grown. To optimize these costs, the next step of the must concentration process can be applied using conventional concentration technology (evaporation) or membrane technology, which has the benefits of a lower energy cost and better quality of concentrate (Jiao et al., 2004). The alternatives for further concentration procedures, shown in Fig. 13.2, are NF, because for NF the osmotic pressure on the permeate side is greater owing to a higher TSS concentration, causing a lower osmotic pressure difference, which results in a better driving force (Kiss et al., 2004; Porter, 1990). The loss of TSS in the permeate of NF should be compensated by returning the permeate of the nanofiltration to the feed of the reverse osmosis, as proposed by Porter (1990) for orange juice concentration. 24
10 9
23
8
Flux, J (kg m–2 h–1)
6
21
5 20
4 3
TSS (%)
22
7
19
2 18
1 0 0
1000
2000
3000
4000
5000 6000 Time (s)
7000
8000
17 9000 10000
Fig. 13.4 Concentration of Furmint must on a pilot-scale reverse osmosis plant in the laboratory of Food Engineering, Corvinus University of Budapest.
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Separation technologies in the processing of fruit juices 391 The second alternative for the production of the final concentrate is membrane distillation. The basic principle of this membrane separation process is the removal of water from the aqueous phase across a hydrophobic membrane when the membrane and the membrane pores are not wetted by the water phase on the feed and permeate side. The driving force of this separation process is the temperature, i.e. the vapor pressure difference on the two sides of the membrane. The separation process can be carried out in the hydrophobic hollow fiber membrane modules, characterized by a large specific mass transfer area of up to 10 000 m2 m–3. Membrane materials include polypropylene, polyvinylidenetetrafluoride, polytetrafluoroethylene (Lagana et al., 2000). The size of micropores is 0.2–1.0 mm. The porosity of the membrane is 60–80% of the volume, and the overall thickness is 80–250 mm, depending on the presence or absence of support. In general, the thinner the membrane and the greater the porosity of the membrane, the greater the flux rate. Experimental results of the concentration of typical Hungarian musts obtained by MD (Kékfrankos and Furmint), and the influence of the type of must, driving force and retentate concentration on permeate flux, are shown in Fig. 13.5. From Fig. 13.5, it is obvious that the influence of the type of must is not significant, but the driving force is the temperature difference, which has the main role in this separation process. Twice the temperature difference resulted in permeate fluxes three times higher, owing to the nonlinear behavior of the vapor pressure in relation to the temperature. OD is a recent membrane process (Lefebvre, 1988) (also known as osmotic evaporation, membrane evaporation, isothermal membrane distillation or gas membrane extraction) which has been successfully applied to the 2.0 Furmint DT 30 °C Kékfrankos DT 30 °C Furmint DT 15 °C Kékfrankos DT 15 °C
1.8
Flux, J (kg m–2 h–1)
1.6 1.4 1.2 1.0 0.8 0.6 0.4 0.2 0.0 10
20
30 40 50 Concentration of retentate CR (TSS%)
60
70
Fig. 13.5 Concentration of typical Hungarian musts Kékfrankos and Furmint, by membrane distillation, influence of the type of the must, driving force and retentate concentration on permeate flux.
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392 Separation, extraction and concentration processes concentration of liquid foods such as milk, fruit and vegetable juice, instant coffee and tea and various non-food aqueous solutions. This technique can be used to extract water from aqueous solutions selectively under atmospheric pressure and at room temperature, thus avoiding the thermal degradation of solutions. It is therefore particularly adapted to the concentration of heat-sensitive products such as fruit juices (Lefebvre, 1988). The process involves the use of a microporous hydrophobic membrane to separate two circulating aqueous solutions at different solute concentrations: a dilute solution and a hypertonic solution (usually salt solution). If the operating pressure is kept below the capillary penetration pressure of liquid into the pores, the membrane cannot be wetted by the solutions. The difference in solute concentrations and, consequently, in water activity of both solutions, generates, at the vapor–liquid interface, a vapor pressure difference causing vapor transfer from the dilute solution towards the stripping solution. The water transport through the membrane can be summarized in three steps: ∑ evaporation of water at the dilute vapor–liquid interface; ∑ diffusional or convective vapor transport through the membrane pore; ∑ condensation of water vapor at the membrane/brine interface. The typical OD process involves the use of a concentrated brine on the downstream side of the membrane as the stripping solution. A number of salts such as MgSO4, CaCl2, and K2HPO4 are suitable. Potassium salts of ortho- and pyrophosphoric acid offer several advantages, including low equivalent weight, high water solubility, steep positive temperature coefficients of solubility and safe use in foods and pharmaceuticals (Courel et al., 2000; Lefebvre, 1988). When compared with RO and MD process, the OD process has a potential advantage which might overcome the drawbacks of RO and MD for concentrating fruit juice, because RO suffers from the limitation of, high osmotic pressure whereas in MD some loss of volatile components and heat degradation may still occur owing to the heat requirement for the feed stream in order to maintain the water vapor pressure gradient. OD, on the other hand, does not suffer from any of the problems mentioned above when operated at room temperature. The most well-known module designed for OD is the Hoechst-Celanese Liqui-Cel membrane contactor with an effective area/ volume of 2930 m2 m–3, a maximum transmembrane differential pressure of 4.08 bar and a temperature operating range of 1–40 °C, and contains microporous polypropylene hollow fibers of Celgard membrane. These fibers are approximately 0.3 mm in external diameter, with a wall thickness of about 0.03 mm; they have a mean pore diameter of about 30 nm and a porosity of about 40%. The experiments of must concentration by osmotic distillation in the laboratories of Corvinus University of Budapest have been carried out using polypropylene hollow fiber modules with CaCl2 as osmotic solution. The
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Separation technologies in the processing of fruit juices 393 details of these experiments are described elesewhere (Kozak, 2005; Rektor et al., 2006). On a laboratory and pilot scale, the final must concentration was over 65 °Brix (Kozak et al., 2008; Rektor et al., 2006). One of the possibile means of reaching a final concentration of must which can be stored at room temperature is the integration of RO and NF. The first step of this concentration technology is concentration by RO but, as shown in Fig. 13.4, it can be used at least up to 23–26% TSS at usual transmembrane pressures (40–50 bar). Assuming that the transmembrane pressure should theoretically be higher in the case of RO and NF, that the experimental data collected at 40–50 bar transmembrane pressures can be used as the basis for the modeling, and that it can be extended to higher pressures, a theoretical design and optimization of a two-step must concentrator was carried out (Kiss et al., 2004; Vatai, 2007). The calculation of the optimal cut off between the RO and NF steps has been carried out by dynamic programming. The basis of these calculations was the total cost of the RO and NF step, as well as the cost of the complete concentration process. For this purpose, it is necessary to develop a relation between the cost and the concentration of the final concentrate. The simplest solution is to find a relation between the yield of the step (Y) and the cost of the concentration process as whole. Assuming, in this first calculation, that the permeate concentrations in comparison with the retentate concentration should be neglected, the equations can be written as:
YRO = (x1 – x0)/x1
where (cp ~ 0)
[13.1]
and in the case of NF:
YNF = (x2 – x1)/x2
where (cp ~ 0)
[13.2]
From the above equations, the relation between the total cost of a step, RO or NF can be written as:
∑ TCOST = ∑ ICOST + ∑ OCOST = f (Y)
[13.3]
where x0 is the feed must concentration in the RO step, x1 is the outlet must concentration from the RO step, and at the same time the inlet concentration of the NF step, and x2 is the final must concentration at the outlet of the NF, and at the same time the final concentration of the must at the outlet of the system of RO–NF concentrator. TCOST is the total cost of the must concentration technology, and for the RO and NF step, respectively, and ICOST is the investment cost of the must concentration technology, and for the RO and NF step, respectively, and OCOST is the operation cost of the must concentration technology, and for the RO and NF step, respectively. Figure 13.6 shows the changes in the total cost (Hungarian forints (HUF)/ year) of the two-step (RO–NF) must concentration plant, and the dependence upon the outlet concentration of the first (RO) [i.e. inlet concentration to
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394 Separation, extraction and concentration processes 1 560 000 1 540 000
Total cost (HUF/year)
1 520 000 1 500 000 1 480 000 1 460 000 1 440 000 1 420 000 1 400 000 1 380 000
15
20
25 30 35 Concentration of retentate (TSS%)
40
45
Fig. 13.6 Changes of the total cost (HUF/year) of the two-step (RO–NF) must concentration plant in relation to the outlet concentration of the first (RO), i.e. inlet concentration, to second (NF) step; using Hungarian prices from 2006 (1 EUR ª 250 HUF).
the second (NF) step], using Hungarian prices from 2006 (1 euro ª 250 HUF) (Vatai, 2007). From this diagram, it can be seen that the minimum system total cost is at 24% TSS concentration, which is the optimum for this two-step system. It means that the must should be concentrated from the feed to 24% TSS in the RO step, and that it then has to be turned to the NF step, when the final concentration of 45% TSS can be reached. These concentration processes should be carried out at theoretical transmembrane pressures of 100 and 110 bar, respectively.
13.5 References Ashurst P. R. (2005): Chemistry and technology of soft drinks and fruit juices, Blackwell Publishing Ltd., UK. Barta J., Körmendy I. (2007): Basic processes of vegetable and fruit processing technologies (In Hungarian), Mezőgazda Kiadó, Budapest. Bélafi-Bakó K., Gubicza L., Mulder M. (2000): Integration of membrane processes into bioconversions, Kluwer Academic/Plenum Publishers, New York, USA. Bélafi-Bakó K. (2002): Membrane separation processes (in Hungarian), Veszprémi Egyetemi Kiadó, Veszprém. Cassano A., Tasselli F., Drioli E. (2007): Ultrafiltration of kiwifruit juice using modified poly(ether ether ketone) hollow fibre membranes, Separation and Purification Technology 57, 94–102. Cassano A., Marchio M., Drioli E. (2006): Clarification of blood orange juice by ultrafiltration: analyses of operating parameters, membrane fouling and juice quality, Desalination 212, 15–27. © Woodhead Publishing Limited, 2010
Separation technologies in the processing of fruit juices 395 Courel M., Dornier M., Herry J. M., Rios G. M., Reynes M. (2000): Effect of operating conditions on water transport during the concentration of sucrose solutions by osmotic distillation. Journal of Membrane Science 170, 281–289. De Carvalho L. M. J., de Castro I. M., Bento da Silva C. A. (2008): A study of retention of sugars in the process of clarification of pineapple juice (Ananas comosus, L. Merril) by micro- and ultra-filtration, Journal of Food Engineering 87, 447–454. El-Bourawi M. S., Ding Z., Ma R., Khayet M. (2006): A framework for better understanding membrane distillation separation process, Journal of Membrane Science 285, 4–29. Fonyó Z., Fábry G. (1998): Basic principles of unit operations (in Hungarian). Nemzeti Tankönyvkiadó, Budapest. He Y., Ji Z., Li S. (2007): Effective clarification of apple juice using membrane filtration without enzyme and pasteurization pretreatment, Separation and Purification Technology 57, 366–373. Hui Y. H., Barta J., Cano M. P., Gusek T. W., Sidhu J. W., Sinha N. Editors (2006): Handbook of fruits and fruit processing, Blackwell Publishing. Jiao B., Cassano A., Drioli E. (2004): Recent advances on membrane processes for the concentration of fruit juices. Journal of Food Engineering 63, 303–324. Kiss I., Vatai G., Bekassy-Molnar E. (2004): Must concentration using membrane technology. Desalination 162, 295–300. Kozák Á., Bánvölgyi S., Vincze I., Kiss I., Békássy-Molnár E., Vatai G. (2008): Comparison of integrated large scale and laboratory scale membrane processes for the production of black currant juice concentrate, Chemical Engineering and Processing 47, 1171–1177. Kozák Á. (2005): Experimental investigation of must concentration by reverse osmosis and membrane distillation (in Hungarian). MSc thesis, Corvinus University of Budapest. Laeson K. W., Lloyd D. R. (1997): Membrane distillation, Journal of Membrane Science 124, 1–25. Lagana F., Barbieri G., Drioli E. (2000): Direct contact membrane distillation: modelling and concentration experiments, Journal of Membrane Science 166, 1–11. Lefebvre M. S. M. (1988): Method of performing osmotic distillation. US Patent 4,781,837, 1 November. Nene S., Kaur S., Sumod K., Joshi B., Raghavaro K. S. M. S. (2002): Membrane distillation for the concentration of raw cane-sugar syrup and membrane clarified juice, Desalination 147, 157–160. Porter M. C. (l990): Handbook of industrial membrane technology, Noyes Data, Park Ridge Rautenbach R. (1997): Membranverfahren, Verlag, Berlin Rektor A., Kozak A., Vatai G., Bekassy-Molnar E. (2007): Pilot plant RO-filtration of grape juice, Separation and Purification Technology 57, 473–475. Rektor A., Pap N., Kókai Z., Szabó R., Vatai G., Bekassy-Molnar E. (2004): Application of membrane filtration methods for must processing and preservation. Desalination 162, 271–277. Rektor A., Vatai G., Békássy-Molnár E. (2006): Multi-step membrane processes for the concentration of grape juice, Desalination 191, 446–453. Thanedgunbaworn R., Jiraratananon R., Nguyen M. H. (2007): Mass and heat transfer analysis in fructose concentration by osmotic distillation process using hollow fibre module, Journal of Food Engineering, 78, 126–135. Vatai G. (2007): Processing of agricultural raw materials and foods by complex membrane separation processes (in Hungarian), Doctoral thesis, Hungarian Academy of Sciences, Budapest.
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396 Separation, extraction and concentration processes
14 Separation technologies in oilseed processing M. A. Williams, Anderson International Corp., USA
Abstract: The two procedures used to separate oil from oilseeds: crushing in mechanical screw presses or extraction with solvent, are described. Optimization of the various preparation steps required for both procedures is discussed. Processing steps for satisfactory recovery of solvent in solvent extraction plants are explored. Some of the currently available equipment for mechanical crushing and for solvent extraction is described. Key words: separation, solvent extraction, oilseeds, screw press.
14.1 Introduction Vegetable oils have been separated from oilseeds for many centuries, starting with relatively simple procedures for easily processed oilseeds such as sesame, peanut, and oil from oil-saturated fruits, such as olives. Gradually, over the years, more sophisticated procedures and better equipment led to the crushing of many oilseeds that did not easily yield their oils. These procedures grew into the unit operations of cleaning to remove dirt and trash, dehulling to remove low-oil/high-fiber components, flaking to provide uniform-thickness particles, and cooking to harden some seed proteins that are too soft to withstand the high pressure exerted by modern screw presses. Some oilseeds, such as cottonseed, require very elaborate preparation for most efficient oil yields. Equipment serving all the unit operations, in particular screw presses, became more durable and more efficient. In addition to mechanical screw pressing, solvent extraction became a method to obtain vegetable oils, and the solvent extraction systems involved their own unit operations and specialized equipment. Today, modern high© Woodhead Publishing Limited, 2010
Separation technologies in oilseed processing 397 pressure mechanical screw press systems and solvent extraction systems operate at high capacity and high efficiency. Small-volume processors prefer mechanical screw pressing to avoid erecting costly solvent extraction systems. High-volume processors prefer solvent extraction to avoid having multiple mechanical screw presses operating in parallel. In this chapter, procedures and equipment used in obtaining oils from oilseeds will be described, mentioning early systems, elaborating on modern systems, and focusing on newer innovations in mechanical screw pressing and solvent extraction.
14.2 Preparation for oilseed processing Typical preparation involves cleaning, dehulling, flaking, and cooking. Some oilseeds require all steps, some only a few steps. 14.2.1 Cleaning Incoming oilseeds are cleaned to remove sticks, pieces of stems and leaves, sand and dirt. Vibrating screens suffice for most oilseeds, but sometimes pneumatic aspiration chambers are also used to separate light-weight impurities from the heavier seeds. Some oilseeds contain appreciable amounts of weed seed, which can be screened out and sold as bird seed. Magnets placed before easily damaged equipment remove tramp iron to protect sensitive equipment from damage. 14.2.2 Dehulling Hulls should normally be removed from oilseeds before screw pressing. Hulls contain low levels of oil. Any hulls entering the screw press with the kernels (or meats) absorb some of the oil released from the kernels, thereby reducing oil yield. Hulls also reduce screw press capacity because a greater volume of material is passing through the screw press. Hulls are high in fiber and are more abrasive, causing greater wear on the processing equipment and greater consumption of horsepower. Dehulling, particularly of soybean, can be done ‘hot’, ‘warm’, or ‘cold’. The soybean seeds are cracked using hammer mills with swinging or fixed hammers or in corrugated breaking rolls, and the cracked seeds are passed through a chamber against a counter-current flow of air, which removes the hulls that have been loosened in the cracking step. Other seeds are dehulled in machines of two basic designs: bar hullers and disk hullers. Larger seeds surrounded by thick, hard shells are dehulled and separated by various methods, sometimes even by hand. Cottonseed, because it comes into the processing plant covered with short cotton fibers (lint), is first delinted using saw delinters, sometimes going through more than one set of delinters. Fibers removed by the different ‘cuts’ find a ready market as cotton batting or in cellulose manufacturing. © Woodhead Publishing Limited, 2010
398 Separation, extraction and concentration processes 14.2.3 Flaking Oilseed kernels or meats are usually cracked into smaller particles and then flattened into thin flakes, if the meats are to be cooked before screw pressing or sent directly to a solvent extractor. This helps to ensure a more uniform cook before screw pressing. Flakes of uniform thickness also provide for a more uniform and more efficient penetration of solvent in solvent extractors. Flaking is usually done with the meats at approximately 10% moisture and a temperature of around 70 °C (158 °F). At that moisture and temperature, the oilseed meats easily smear into flat flakes without shattering into fine particles. Two types of flakers are used. One type employs two heavy steel rolls in parallel, almost touching each other, and revolving in opposite direction, so that both rolls pull the meats down between them. Mechanical jack screws or hydraulic cylinders apply pressure against the bearing blocks supporting the rolls so that the rolls do not spread apart when the meats pass between them. The second type consists of a set of flaking rolls stacked on top of each other, usually five high. This has an advantage in that the weight of the upper rolls provides pressure on the lower rolls to make thin flakes. The meats flow down the stack of rolls and are progressively flattened into thinner and thinner flakes. 14.2.4 Cooking Many oilseeds have soft meats, which are even softer when the hulls are removed. Such dehulled meats can easily form an oily paste within the screw press. These meats are usually cooked to stiffen and harden the protein so it can better withstand the pressures generated by the screw press. Cooking is also used to deal with undesired heat-sensitive components such as undesirable enzymes in some oilseeds and with gossypol in cottonseed. Gossypol can impart a red color to the oil and render the deoiled solids unfit for some animal feeds. Cooking can combine the gossypol with the protein in the solids preventing the gossypol from being extracted with the oil. If the combination is strong enough, it will reduce the gossypol’s toxic effect in defatted solids that go into animal feeds. Two types of cookers are used to cook oilseeds: ‘stack cookers’ and ‘horizontal cookers’. A stack cooker is comprised of steam-heated chambers stacked in a vertical housing. Steam jackets on the bottom surface of each stack supplies the heat for cooking. A horizontal cooker consists of a bank of two or more cylindrical vessels positioned horizontally, usually one above the other. Both horizontal vessels have the outer walls steam jacketed. An advantage of horizontal cookers is that raw feed drops into a large volume of tumbling material that is already hot; thereby quickly coming up to desired cooking temperature. In a stack cooker, on the other hand, the raw feed falls on top of a thick layer of slowly agitated material that is agitated
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Separation technologies in oilseed processing 399 beneath the surface, causing some delay in raw material reaching desired cooking temperature. If the raw material contains an enzyme that damages oil quality, the enzyme can become active long enough inside the top stack to do its damage before the temperature becomes high enough to inactivate the enzyme. This long delay in reaching the desired cooking temperature does not happen in a horizontal cooker. There is a slight possibility, however, that some undercooked oilseed could move across the surface of the oilseed in the cooker instead of being churned down into lower levels within the cooker. If this happens, then some undercooked oilseed could move quickly toward the discharge port, leaving the cooker before it is fully cooked. If an enzyme is present, and if it is critical that the entire oilseed is exposed to adequate retention time, it would be prudent to use two horizontal cooking vessels in series before sending the oilseed on toward downstream equipment. Cooking is done at elevated moisture, usually 9 to 12%, which is too high a moisture level for efficient pressing. Cookers are followed by similar vessels that serve as dryers to bring the moisture down to 2–5%. This is done in vented vessels, usually with a draft of air passing through them, and with a heat source, usually steam jackets, to elevate the oilseed temperature high enough to drive the moisture out, usually 104–120 °C (219–248 °F).
14.3 Extrusion preparation for oilseed processing 14.3.1 Origin of extrusion Extrusion cooking expanders were accepted into pet food manufacture in 1954. These machines cooked and puffed cereal grains and nutritionally balanced feed formulations for the petfood industry. In the 1960s, cooking extruders were applied to the preparation of rice bran ahead of solvent extraction. In the 1970s, extrusion was also used to form porous collets from flaked soybean and other oilseeds. Extrusion brings the incoming oilseed to desired moisture, temperature, and pressure conditions within 10 s. If there are troublesome enzymes present, extrusion can bring the oilseed enzymes to inactivation conditions long before the enzymes have time to do damage. Extrusion is even more effective to inactivate enzymes than the horizontal, atmospheric pressure cookers described earlier. Extrusion has been successfully employed for inactivating enzymes such as lipase in rice bran (Williams and Baer, 1965) and urease in soybean (Williams, 1991). 14.3.2 Extrusion before solvent extraction The first application for extrusion before solvent extraction was rice bran (Baer, 1966). Rice bran arrives at the plant as a fine powder, which presents two problems: the fine powder is very difficult to process because solvent © Woodhead Publishing Limited, 2010
400 Separation, extraction and concentration processes cannot percolate through a bed of fine powder. Even more troublesome, the bran contains the enzyme, lipase, which splits vegetable oils rapidly into free fatty acids. Activated as soon as the bran is exposed to air, lipase will raise the free fatty acid level by approximately 3 to 7% every day until the level reaches 50–75%. Cooking conditions during extrusion quickly inactivates lipase and converts the bran into a coarse meal that permits rapid percolation of solvent through the bed of material within the extractor. Similar extrusion also transforms flaked soybean and flaked cottonseed into porous collets that handle better in a solvent extractor. Farnsworth described extrusion of cottonseed (Farnsworth et al., 1986). 14.3.3 Closed-wall extruders An extruder consists of a rotating wormshaft within a cylindrical barrel (Fig. 14.1). Material to be extruded is fed into one end of the barrel, and the wormshaft forces it out through a die plate at the discharge end. The wormshaft flighting is not a continuous wrap; it is composed of partial wraps and unflighted segments. Stationary pins protruding from the barrel wall intermesh between the individual wraps. The combination of rotating flights and stationary pins masticates the oilseed, quickly blending injected steam into the oilseed to elevate the moisture and temperature within the extruder.
Fig. 14.1 Typical closed-wall extruder (courtesy Anderson International Corp.).
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Separation technologies in oilseed processing 401 The oilseed within the extruder is subjected to high pressure, 1379–4137 kPa (200–600 psi), and is pressure-cooked at optimum moisture and temperature to convert the protein into a tacky, elastic‑like condition. The injected steam releases its heat of vaporization, which helps to heat the oilseed. Additional heat is generated by friction from the rotating shaft. 14.3.4 Slotted-wall extruders Oilseeds containing more oil than cottonseed meats (30–33%) are difficult to process through a closed-wall extruder because the liberated oil accumulates within the extruder and disturbs the steady‑state operation. Also, the extruded collets cannot reabsorb all the liberated oil, and some oil is lost in product transport. If this is troublesome in extrusion plants, sometimes deoiled meal is blended with the incoming oilseed to dilute the oil level. This corrects the problem of excess liberated oil, but the recycling of meal increases the work done by the extruder, the solvent extractor, and the meal desolventizer, and reduces the total overall capacity of the oil mill. Anderson International (Anderson) introduced a slotted-wall extruder that permits a controlled release of excess oil through the slotted wall and produces collets at 20–30% oil (Williams, 1990) (Fig. 14.2). This slotted-wall extruder can process full fat safflower (at 42% oil), sunflower (at 42–44% oil), and peanut (at 45% oil) producing collets at 20–30% oil for solvent extraction. The liberated oil comes through the slotted wall and thereby
Fig. 14.2 Typical slotted-wall extruder (courtesy Anderson International Corp.).
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402 Separation, extraction and concentration processes bypasses the solvent extractor. Typical preparation before extrusion is to crack to approximately 1.6 mm (1/16≤) particles or flake to 0.3 mm (0.012≤); heat to approximately 60–71.1 °C (140–160 °F); and reduce moisture to approximately 8% without cooking the protein. Closed-wall and slotted-wall extruders are also fitted with discharge cone chokes rather than die plates at the discharge end. The cone is mushroom shaped and is moved laterally toward and away from a matching socket mounted on the extruder’s discharge. A hydraulic cylinder moves the cone in and out. Pressure generated by the extruder causes the oilseed to bear against the cone. Hydraulic fluid trapped within the cylinder holds the cone in a fixed position. Extruded product flowing across the cone into atmospheric conditions ‘expands’ with internal pores similar to the way collets expand when exiting the dies in a die plate. The sheets of porous cake falling past the cone break up in transit similarly to the way collets break up. 14.3.5 Dry extrusion Extrusion is also used to prepare oilseeds for subsequent screw pressing. Nelson et al. (1987) processed coarsely ground soybean through a high shear, closed-wall extruder, thus obtaining a hot, foamy mixture of small particles, liberated oil, and vaporizing moisture. The product upon discharge acted and looked like a fluid owing to moisture boiling through the freshly liberated oil. The soybean was introduced into the extruder at ambient temperature and moisture. The extruded product exited at 135 °C (275 °F), and the excess heat caused the moisture to flash down to 6–7% moisture. The heat is generated by friction through a specific configuration of screws, sometimes including steam locks, and by a cone attached to the end of the rotating shaft positioned into a stationary conical socket attached to the discharge plate. The contour of the socket matches the rotating cone and has a circular orifice in the center of the socket through which the extrudate exits. Shear is influenced by the proximity of the rotating cone to the stationary socket. 14.3.6 Dox-Hivex™ extruders Extruders adapted for dry extrusion, both closed‑wall and slotted‑wall, can be fitted with adjustable jaws or a manually positioned plunger to apply a choking action against the discharging oilseed (Fig. 14.3). These extruders are usually operated at low moisture so as to provide for high-shear rupturing of the oil cells. With little or no preparation, oilseeds are converted to a foamy, semi‑fluid extrudate at 121–148.9 °C (250–300 °F) that, as it cools, flashes down to 5–7% moisture. Shear is also influenced by the proximity of the cone point to the discharge port and the jaws and by positioning the jaws relative to each other or, if a plunger is used, by positioning the plunger relative to the discharge port. The plunger, a recent improvement replacing
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Separation technologies in oilseed processing 403
Fig. 14.3 Dox-Hivex™ extruder (courtesy Anderson International Corp.).
the jaws with a manually positioned plunger, can be moved close enough to the discharge port to almost completely block any discharge of material. 14.3.7 New developments in extrusion Extrusion preparation ahead of full-pressing has proven to be much more effective than traditional preparation. Two screw press manufacturing companies (that also manufacture extruders), Anderson International Corp and Insta-Pro, have developed extrusion preparation procedures pioneered by Nelson et al. (1987).
14.4 Mechanical pressing of oilseeds A mechanical screw press (Fig. 14.4) accepts a continuous stream of oilbearing material, compresses it under very high pressure exerted by a wormshaft rotating within a slotted-wall cage, the flights on the wormshaft propelling the oilseed forward as the shaft exerts pressure. The pressure releases the oil, most of which flows out through the slotted-wall cage. The deoiled solids flow through an adjustable port, like a movable cone, at the discharge end of the press (Fig. 14.5). The most common screw press application is full-pressing. Oilseeds are processed to liberate as much oil as possible through the full-press. Residual oil levels are usually measured as ether-extract using AOCS
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404 Separation, extraction and concentration processes
Fig. 14.4 Typical screw press (courtesy Anderson International Corp.).
Fig. 14.5 Cone choke on screw press (courtesy Anderson International Corp.).
method Ba 3-38 (or other equivalent and industry accepted procedures). A full-press generates an internal pressure of around 96 500 kPa (14 000 lbs in–2) so as to press out as much oil as possible. Earlier small-capacity full-presses reduced residual oil levels to between 3 and 5% in the exiting solids. Currently, higher capacity full-presses produce residual oil levels of 5–8%. With minor modifications, the same presses are (and always have been) used to partially press oilseeds as preparation in front of subsequent
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Separation technologies in oilseed processing 405 full-presses or as preparation for solvent extraction. The objective in prepressing is to arrive at 15–25% residual oil and run at higher capacities consuming less horsepower. Oil liberated by a screw press contains around 2–5% solids, but can contain up to 15% solids. The solids are removed in two steps. Step one employs a rectangular settling chamber equipped with a drag assembly that drags settled solids from the floor of the chamber, up an end wall, and over a wedge-wire drainage screen so that the solids can free drain as much as possible before recombining with fresh material going to the screw press. Step two clarifies the oil in a horizontal pressure leaf filter. Centrifuges, especially if they are three-phase (introducing water to flush solids out of the oil), are not well received in oilseed press plants because centrifuged solids, if reintroduced into the feed stream, are more difficult to press than solids obtained from filter presses. The filter press solids drop into a hopper equipped with a variable speed discharge screw that meters a stream of filter cake to blend in with fresh material going to the press. The feed to the screw press should be a homogeneous blend of fresh material, drainage chamber solids, and filter press cake. If the ratios of the three components are kept constant, the screw press will operate in a steady-state mode. 14.4.1 New developments in screw pressing There is a growing interest, today, in obtaining vegetable oils for food uses without the use of chemical solvents. The long-standing traditional method of achieving this is screw pressing; and screw pressing has been in use for more than a century. Since their introduction, screw press requirements have changed considerably from a time when labor was inexpensive and capital equipment costs were of concern. Today, labor costs are very high, and any simplification of time-consuming maintenance procedures would, of course, make mechanical screw presses much more attractive. A case in point is the tedious process of putting new barrel bars and spacing clips into the drainage cages. Historically, drainage cages are constructed of two cage halves that are clamped tightly together to surround the worm shaft propelling the oilseed through the screw press (Fig. 14.6). Barrels bars are assembled into the cage half, by placing a cage half horizontally on a work stand and manually positioning barrel bars and thin metal spacing clips of various thicknesses (Fig. 14.7). The spacing clips are used to provide gaps between the bars where oil can escape from the cage. Because very high pressure exists within the cage, the barrel bars must be captured tightly within the supporting structure of the cages to prevent any shifting of the bars during operation. This is a very time-consuming and labor-intensive task. However, if the assembly is not done correctly, the screw press will perform poorly. A desirable improvement is for screw press manufacturers to provide cages
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406 Separation, extraction and concentration processes
Fig. 14.6 Set of cages and shaft (courtesy Anderson International Corp.).
Fig. 14.7 Placing spacers (courtesy Anderson International Corp.).
that are easier to assemble and do not require so much precision in assembly. All screw press manufacturers are looking at this. The cylindrical drainage cage is constructed in two longitudinal halves adjacent to the wormshaft, sections are bolted tightly together to withstand © Woodhead Publishing Limited, 2010
Separation technologies in oilseed processing 407 the pressure. Inserted between the two cage sections are two lug bars called ‘knife bars’, 180° apart, running the entire length of the cage. Both knife bars have lugs protruding through the channel and intermeshing with the individual worm flights. The shaft is either of one piece construction or is made of individual worm segments slid onto a central shaft and keyed into position. Wormshafts are configured to progressively increase compaction on the oilseed as it is pushed from worm segment to worm segment. This ensures that compaction does not diminish as the volume of the oilseed is reduced owing to both the compression of the solids and the escape of noncompressible oil. This is done by reducing channel depth (the open space between the inside diameter of the barrel and the hub surface of the central shaft) and by decreasing the pitch of successive worm flights. Sometimes both techniques are employed. Some screw presses employ a force feeder to ensure sufficient inlet capacity to achieve desired screw press capacity. This introduces a second wormshaft and barrel plus a second drive mechanism to power the force feeder. The force feeders on large capacity screw presses are always driven by a smaller horsepower motor than the main wormshaft. Sometimes the smaller drive motor of the forced feeder reaches full horsepower consumption when the main drive motor is partially loaded. More input cannot be introduced because the force feeder will overload, but the main drive motor could easily accommodate greater capacity. Some screw press manufacturers are designing screw presses with a single motor driving both shafts or a gravity fed inlet not requiring a separate drive. A single drive motor allows for maximum available horsepower consumption on the main motor and does away with the problem of balancing the motor loads between the two drive motors. Some screw press manufacturers are working on new technology to improve capacity, reduce horsepower consumption and, by reducing horsepower consumption, reduce wear. Studies are also underway to characterize by mathematical modeling how a wormshaft configuration (profiles of pitch, channel depth, and thickness of flighting) influences the progressive compaction of the oilseed as it is conveyed by the wormshaft through the drainage cage. Most screw press manufacturers have shaft profiles that they know from experience have given satisfactory performance in the past. Currently, with the requirement of high capacity production plants, especially with concern of producing products that have not been exposed to chemicals such as hexane, there is interest in very high capacity screw presses that require less connected horsepower. Target capacities of 150 Mt d–1, or more, pressing to 3–5% residual oil using a smaller than traditional size motor is desirable, but no full-press offered today can achieve that performance. Screw press manufacturers are looking at mathematical modeling of shaft profiles to assist in more optimum shaft design. The objective is to design shaft profiles that press out maximum oil and dissipate minimum energy into friction so as to achieve maximum capacity at minimum horsepower consumption.
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408 Separation, extraction and concentration processes Most manufacturers are not going to give away their trade secrets, but some of the results of these efforts have already been published. Anderson International published some of their work in 1983 (Williams, 1983). Vadke et al. (1988) of the University of Saskatoon, developed a mathematical model of screw press operation by superimposing filtration analysis onto screw press theory. Predictions of how changes in shaft rotation and choke opening would affect screw press performance obtained by the mathematical model agreed reasonably well with experimental results obtained using a small laboratory screw press. Screw press oil yields can be increased by using supercritical CO2 to assist in obtaining the oil from a traditional screw press. Experiments in the early 1980s showed that significant increase in oil yields and in oil quality can be achieved by using supercritical CO2 (Friedrich et al., 1982) and (Friedrich and Pryde, 1984). Supercritical CO2 injected into the oilseed within the screw press cage can leech through the oilseed, dissolving the oil, and flow out with the oil through the drainage cage. Voges et al. (2008), reported oil yields increasing from 27% to as much as 71% when supercritical CO2 is used. A joint venture between two manufacturers is offering systems to do this (14.4.2). 14.4.2 Full-press suppliers There are many suppliers of small capacity screw presses offered for niche markets of special oils and small biodiesel plants. Vegetable oil production of oils for food purposes, however, is generally done in high capacity production plants where high capacity screw presses are desired. A list of well-known suppliers of high capacity screw presses and solvent extractors is available (Anon, 2009b). These suppliers also offer auxiliary equipment, and some of them offer entire plants. Anderson International has offered screw presses for many years. Their presses have been very well received in the past, but newer demands for screw presses favor higher capacity, less horsepower consumption, and simpler maintenance. Anderson has been working on a full-press of high capacity, consisting of a single motorized wormshaft and containing a simpler, less expensive choke (Fig. 14.5). One model, the Victor600™ (Fig. 14.8), is currently in field operation; another larger model is planned. The Victor600™ having a 6≤ (152 mm) diameter bore (inside barrel diameter) is rated at 50–75 Mt d–1 capacity, and is operating on sunflower and canola. The Victor1200™ with a 12≤ (304 mm) diameter bore is rated for 150 Mt d–1. Both presses feature a new choke design that permits replacing worn choke parts without disassembly of the wormshaft or removal of the drainage cages. Anderson also manufactures extruders and has developed an extrusion preparation system employing an extruder with a drainage cage followed by a full-press to press out the remaining oil. This ‘combo’ system (Fig. 14.9), utilizes a slotted-wall extruder for preparation, which can significantly increase the capacity of the full-press.
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Separation technologies in oilseed processing 409
Fig. 14.8 Victor 600™ press (Courtesy Anderson International Corp.).
Desmet Ballestra offers their line of Sterling Series presses (Fig. 14.10). These screw presses have 8≤ (203 mm) bores up to 13.75≤ (349 mm). Desmet purchased Rosedown several years ago, and some of the Sterling presses are still referred to as Rosedown Presses (www.rosedowns.co.uk). Sterling presses feature ‘multi-stage’, low-compression worm assemblies, and have an internal cake breaker. They can also be equipped with a ‘barring motor’, which is used to empty the press after power outages without requiring opening of the cages. Desmet is a major supplier of biodiesel plants around the world and has partnered with OriginOil (http://www.originoil.com/about-us/company.html) of Los Angeles, CA to commercialize OriginOil’s patent-pending process to extract oil from algae (http://www.originoil.com/company-news/originoilfiles-international-patent-for-low-energy-high-efficiency-algae-production. html). OriginOil’s patent is for a method to grow algae and provide sufficient light and CO2 for rapid growth, a system for rupturing oil cells, and for producing electricity with closed-loop CO2 recycling. The new process does not require drying the algae, which provides a substantial savings in electrical and thermal energy costs over traditional processing http://www. originoil.com/company-news/originoil-announces-partnership-agreementwith-desmet-ballestra-at-naa-conference.html. Dupps makes five models of Pressor® with capacities of 1500 to 12 000 lbs h–1 (18–144 Mt d–1). Bore diameters are 7≤ to 13≤ (177–330 mm) (Fig. 14.11). Pressors have hydraulically operated chokes that can automatically exert uniform discharge pressure by means of a controlled hydraulic pressure maintained on the hydraulic cylinder. The choke plunger ‘floats’ in equilibrium between the dynamic forces generated by the full-press and the force exerted by the hydraulic cylinder. When hydraulic pressure is held constant, the © Woodhead Publishing Limited, 2010
410 Separation, extraction and concentration processes Cottonseed Steam
Conditioner
Steam vapors
200 TPD Cottonseed meats
Condensate
Dox/Hivex™ extruder
Steam vapors
Expanded meal 175 TPD Vegetable oil to filtration 25 TPD
Steam
Dryer
Condensate
Expanded meal 167.4 TPD
11D expeller®
Total oil to filtration Hivex: 25 TPD Expeller: 26.6 TPD Total: 51.6 TPD
Vegetable oil to filtration 26.6 TPD
Press cake at 6% R.O. to meal grinding 140.8 TPD
Fig. 14.9 Extrusion–screw press system (courtesy Anderson International Corp.). TPD, metric tons per day; RO, residual oil.
plunger ‘floats’ in and out to ensure that the pressure within the screw press remains as steady as the oil pressure within the hydraulic cylinder, which pressure being controlled, does remains steady. © Woodhead Publishing Limited, 2010
Separation technologies in oilseed processing 411
Fig. 14.10 Desmet Ballestra press (courtesy Desmet Ballestra).
Fig. 14.11 Dupps screw press (courtesy The Dupps Company).
French Oil Machinery (Fig. 14.12) offers a line of Achiever presses; six models for full-pressing, which can also be used for pre-pressing before solvent extraction. Capacities range from 11 to 136 Mt d–1 on full-pressing. Achiever presses have force feeders, water-cooled drainage cages as well as water-cooled main worm shafts, and a cone choke to create back pressure against forward flow of material through the press. Harburg-Freudenberger offers several models of presses for full-pressing and pre-pressing. The largest full-press is Model EP-22 (Fig. 14.13); rated for 100–120 Mt d–1 and equipped with a 250–400 kW (339–543 hp) motor. Residual oils are 5–8%. Choking is accomplished through a fixed throttle
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412 Separation, extraction and concentration processes
Fig. 14.12 French OM press (courtesy The French Oil Machinery Company).
Fig. 14.13 Harburg-Freudenberger press (courtesy Harburg-Freudenberger).
ring that replaces mechanically adjustable chokes. Harburg-Freudenberger has also developed a worm design based on pressure worms to build up pressure alternating with pressure equalizing and relaxation via internal conical throttle rings. Harburg-Freudenberger has entered into a joint venture agreement with Crown Iron Works to offer Harburg-Freudenberger screw presses employing liquid CO2 injection to improve release of oil (the Hiplex™ process). Screw presses using this technology can reduce residual oils from 5–8% to 3–5%. Safe Soy Technologies, Elsworth, Iowa, recently put a Hiplex™ system on stream on soybean (Radio Iowa, 2008). Experimental work elsewhere in Hamburg University of Technology, reports significant increase in oil yields from soybean and rapeseed by injecting CO2 into a screw press (Voges et al., 2008). Willems P, Kuipers N and de Haan A, of University of Twente,
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Separation technologies in oilseed processing 413 working with sesame, linseed, rapeseed, palm kernel, and jatropha, reported oil yields 30% improved over conventional screw pressing (Willems, 2007), and claimed that displacement of oil by dissolved CO2 was the major cause of increased oil yields. The CO2 can be recovered, but recovery is expensive, and sometimes, in smaller capacity plants, the value of the extra oil does not justify the added cost to recover CO2. Insta-Pro International offers several models of extruders for preparation of oilseeds before pressing, the two largest extruders being Model 9400 with 300 connected horsepower (224 kW), and the Model 2000 Double Barrel with two 150 hp (112 kW) motors (Fig. 14.14), both models capable of 6720–7680 lbs h–1 (3056–3490 kg h–1) capacity. The extruded oilseed can then pass into several models of Insta-Pro screw presses, the largest being Model 5005, with a 60 hp (44 kW) motor, which can achieve 4000–4400 lbs h–1 (1818–2000 kg h–1) capacity (Fig. 14.15). 14.4.3 Other presses Duyvis Wiener BV has been manufacturing continuous pot presses for obtaining cocoa butter from cocoa bean (Fig. 14.16). The cocoa beans are ground/liquefied into a slurry (liquor) and heated, to ensure that all the butter is melted and a pumpable mass is produced. The liquor is pumped at a temperature of around 100 °C (212 °F) into the press until the press is fully charged. Pumping then stops and a hydraulic ram applies pressure to squeeze out the melted butter to 10–12% residual butter in the cake. Duyvis
Fig. 14.14 Insta-Pro extruder (courtesy Insta-Pro International).
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414 Separation, extraction and concentration processes
Fig. 14.15 Insta-Pro screw press (courtesy Insta-Pro International).
Fig. 14.16 Duvis press (courtesy Duvis Wiener BV).
has adapted this very much different type of press for pressing slurries of oilseeds other than cocoa bean, like peanuts and sesame seeds, with good results. The oilseeds are first ground to a slurry then heated to around 60 °C (140 °F). The heated slurry is then pumped into pots, and the pump pressure forces the vegetable oil to escape through fine-mesh grids built into the pots. The pump can exert a pressure of up to 2000 kPa (290 psi). Pumping then stops and a hydraulic ram forces the pots, with their contents, tightly together, to press out more oil. Final residual oils of 8–10% can be reached, but at © Woodhead Publishing Limited, 2010
Separation technologies in oilseed processing 415 the expense of capacity. Pressing cocoa to 22–24% residual oil can be done at 3290 kg h–1 (78 Mt d–1). Pressing to 10–12% residual oil can be done at capacities of 1315 kg h–1 (31 Mt d–1). The press is then opened, the solids are pushed into a conveyor, and the press closes for a new cycle. Cycling can be accomplished manually or automatically through a microprocessor.
14.5 Percolation solvent extraction in oilseed processing Solvent extraction is a very effective method for recovery of oil, especially with materials low in oil. Full-pressing can reduce oil content to a minimum of 3–5%. This is fine for oilseeds containing large amounts of oil, but for oilseeds low in oil (such as soybean at 18%), a reduction in oil from 18% to 3–5% leaves a reasonable amount of the available oil remaining in the presscake (approximately 16–27% of the total oil content). For that reason, solvent extraction is preferred for oilseeds of low oil content if capacities are sufficiently high to justify the expense of a solvent-extraction plant. Solvent extractors operate at lower temperature than full-presses, so the oil can be of higher quality if any components are heat sensitive. However, solvents do extract some components that are not true oils (triglycerides), thus necessitating some downstream procedures for removing them. Also, consumers are becoming reluctant to eat food products that have been in contact with chemical solvents. Percolation extractors work on the same principle as coffee percolators. Heated solvent is rained through the prepared oilseed, dissolving out the oil. Unlike a coffee percolator, however, the oilseed is extracted several times with the solvent redirected through the oilseed in a countercurrent flow pattern. Fresh solvent passes through the oilseed after it has been exposed to multiple passes of solvent. The fresh oilseed, on the other hand, as soon as it enters the extractor is washed with the solvent, now containing a reasonable amount of oil, just before the solvent/oil mixture (miscella) leaves the extractor. To ensure that the miscella contains a minimum of solids, most percolation extractors direct the final miscella one more time through the bed of solids (after the solids have had a chance to settle down into a firmer bed) in order to use the bed of solids as a filter to remove as many solid particles as possible. The extracted solids, after the final wash with incoming fresh hot solvent, remain in the extraction vessel to allow the solvent to free drain as much as possible before the finished extracted solids (marc) are removed from the extraction vessel. The marc then passes into a desolventization vessel. The miscella goes on to an evaporator followed by a stripper to remove all traces of solvent from the oil. Most percolation extractors permit external control of where the miscella from each stage goes. Valves downstream from each sump pump can send some miscella back to the stage from which it came and, by using different valves and different piping, send some to the preceding stage. Sometimes © Woodhead Publishing Limited, 2010
416 Separation, extraction and concentration processes all the miscella is pumped back over the stage from which it came. This allows the miscella to make multiple passes within each stage to increase solvent contact without requiring greater input of fresh solvent and also helps to maintain an adequate head of miscella above the bed of oilseed. The counter‑current movement of miscella from stage to stage, when the above is done, occurs through overflow weirs built into each sump. Oilseed preparation, especially particle‑size and moisture, influence how fast solvent can percolate through the bed of solids. Percolation should be rapid enough to ensure good contact of every particle with solvent. If percolation rate is too slow, the solvent will flood over the surface of the solids, greatly impairing the performance of the extractor. Similarly, if percolation is so rapid that the solvent flows down through the center of the bed without spreading across the entire bed, some of the oilseed will not come into contact with the solvent. Marc from flaked oilseeds usually retains 30 to 40% by weight of solvent. Pre-pressed cake and extrusion prepared oilseeds can free drain to 20–25% solvent before leaving the extraction vessel. This lower retention of solvent by the marc is one of the major benefits of extrusion before solvent extraction. If the marc free drains to a low solvent level, substantial energy can be saved in desolventization because less solvent has to be driven off by thermal evaporation. However, the solvent that drains out of the marc ends up in the miscella, which increases the solvent that has to be evaporated from the miscella. In all instances, where extrusion preparation is used, the solvent flow into the extractor can be lowered, resulting in less solvent in the miscella as well as less solvent in the marc. Extruded collets permit this reduction in solvent-to-meal ratio because extrusion helps release the oil from the solids, enabling most of the oil to extract very quickly. Rapid percolation also helps by flushing out the released oil almost as soon as the solids enter the extractor. The more deeply embedded oil still needs time to be fully extracted, but less solvent is required, thus permitting an extraction plant to reduce energy consumption for both miscella and marc desolventization. 14.5.1 Rotary extractors Rotary extractors have a rotating circular disk-shaped bed of oilseed slowly spinning within the extraction chamber. These extractors are quite large, and the bed depth of oilseed can be 2–4 m. Fresh hot solvent is directed over the bed at the end of the extraction cycle, just before the marc is allowed to free drain. The oilseed is extracted in a counter-current fashion as described above. Figure 14.17 shows a schematic of a typical rotary extractor. 14.5.2 Perforated belt extractor Another type of percolation extractor uses a longitudinal, perforated belt carrying an extended bed of oilseed. The belt forms a flexible endless loop
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Separation technologies in oilseed processing 417 Top and bottom bearings accessible from outside the unit for easy maintenance
Slurry filling spout
Ultra-free turning spindle
Sealed dividers form baskets to ensure miscella stage separation
Conical top with sight glasses for maximum visibility
Reliable bevel gear drive Self-cleaning screen for outstanding drainage Miscella collection pan
Sealed dump hopper to prevent contamination
Fig. 14.17 Schematic of rotary extractor (courtesy Desmet Ballestra).
moving over pulleys at both ends and carrying the oilseed on the top surface through the extraction vessel. The belt is on a slight incline, and the solids travel up the incline toward the discharge end. This helps to ensure a more efficient countercurrent flow of solids and miscella while in the extractor. The marc falls off the higher end as the belt passes around the discharge pulley. Fresh oilseed is dropped onto the lower end of the belt after it comes to the top of the feed end pulley. 14.5.3 Sliding cell extractor A sliding cell extractor slides the solids along a stationary steel plate that is perforated to allow miscella to drain through the plate while the bed of solids
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418 Separation, extraction and concentration processes slide across the surface of the plate. The incoming oilseed falls onto an endless belt mechanism consisting of tall vertical plates that form closed receptacles or cells into which the oilseed is captured. Countercurrent miscella washes are introduced onto the bed of cells as they pass under miscella spray heads. The moving cell assembly wraps around a drive pulley, turning the oilseed contents upside-down and sliding them across a second perforated plate as the cells return back to the feed end of the extractor. Each cell empties just before the moving assembly wraps around a second pulley. The now empty cells then move under a hopper from which fresh oilseed enters the cells. 14.5.4 Rectangular loop extractor A rectangular loop extractor drags the oilseed through a chamber shaped like a closed loop. This extractor uses an ‘en masse’ conveyor that moves the oilseed through a closed housing with a rectangular cross section. The housing loops back on itself, looking somewhat like a collapsed doughnut resting on its side. Fresh oilseed is conveyed through an inlet onto the upper level and is sprayed with rich miscella. The oilseed encounters multistage counter-current extraction as it travels through the loop. The final wash is with fresh solvent followed by a free drain period before the extracted oilseed exits. Figure 14.18 shows a rectangular loop extractor. 14.5.5 Extractor suppliers Major suppliers of solvent extraction equipment are Crown and Desmet. Other suppliers are Harburg-Freudenberger and Lurgi GmbH and several others around the world. A list of suppliers of oilseed processing equipment is available (Anon, 2009b). Many of these companies also supply auxiliary equipment and can supply entire plants. Crown specializes in rectangular loop extractors (Fig. 14.18). Crown offers their Model 3 extractor, which is rated for 9000–9500 Mt d–1 on flaked oilseeds and up to 10 000 Mt d–1 on extruded collets. Crown has made advances in hexane recovery, especially in stripping and desolventation. Their Super Stripper System™ can result in significantly lower solvent losses than the 3.78 L t–1 (1 gal ton–1) that was common 10 years ago. If the equipment is operated correctly and kept in good repair, solvent losses of 0.8–1.9 L Mt–1 (0.22–0.5 gal ton–1) can be achieved. Desmet makes two types of solvent extraction vessels: the Reflex® solvent extractor, a rotary extractor, and the LM™ extractor, a perforated belt extractor. Capacities of Desmet extractors range from 500 to 10 000 Mt d–1 on flaked soybean. Desmet’s Reflex® rotary extractor has a horizontal circular extraction chamber divided into sections like a cut pie. The material to be extracted revolves under a stationary solvent/miscella spray harness and feed input chute, which directs the incoming oilseed and solvent/miscella over the appropriate sections of oilseed as the bed of material rotates under
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Solids inlet hopper with electronic level sensor
Extractor drive speed controlled by the incoming volume of raw solids read by the electronic sensor
First wash
Hydroclone
Hydroclone miscella clarifier Full miscella outlet
Countercurrent recycle stages
The Crown Hydroclone removes the final traces of fines from the full miscella. The miscella can be pumped directly to the evaporation system.
Fresh solvent rinse
Drainage section
Final recycle
Bar screen Self-cleaning Countercurrent stationary vee-bar recycle stages screens
The flake bed acts as a brush - it continually shears the stationary bar screen clean of flow-obstructing fines
Fig. 14.18 Rectangular loop schematic (courtesy Crown Iron Works Company).
Separation technologies in oilseed processing 419
© Woodhead Publishing Limited, 2010 Extracted solids outlet
420 Separation, extraction and concentration processes the spray harness. The oilseed within the pie-shaped sections slides over a stationary perforated bottom. The sliding of the oilseed over the bottom helps to keep the perforations open. When extraction time is finished, the extracted oilseed (marc) slides over a lower stationary marc hopper, and the marc drops into the hopper. The rotating extraction chamber then moves under the stationary feed inlet hopper, which drops fresh oilseed onto the revolving disk. Desmet also supplies perforated belt extractors. The belt is made of a series of wide, rectangular sections of either wire mesh or wedge-bar screens that are connected together to form a looped endless-conveyor belt. Figure 14.19 shows the Desmet LM™ extractor, and Fig. 14.20 shows a schematic of this extractor. The perforated belt extractor has only one moving part: the belt. Belt speed is automatically adjusted to maintain a constant level in the inlet hopper. The belt is self-tensioning, self-cleaning, and fully automatic. The extractor is designed to permit various extraction and drainage times before the marc discharges from the extractor. Typical times for soybean are 42 min for extraction and 15 min for drainage. Other duration times are set for different oilseeds. Harburg offers their carousel extractor, which is a rotary extractor with the chamber containing the oilseed rotating beneath a stationary spray head/input hopper assembly and above stationary miscella sumps and marc hopper assembly. A carousel extractor can be supplied double-decked, and its capacity is 50 to 4000 Mt d–1. Lurgi offers a line of sliding cell extractors with a capacity range of 100
Fig. 14.19 Perforated belt extractor (courtesy Desmet Ballestra).
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1
2
3
2
Fig. 14.20 Perforated belt schematic (courtesy Desmet Ballestra). 1, Fresh oilseed input; 2, fresh solvent input; 3, miscella output; 4, marc output.
Separation technologies in oilseed processing 421
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422 Separation, extraction and concentration processes to 4000 Mt d–1. The Lurgi sliding cell extractor provides dual-pass, shallowbed extraction with the partially extracted oilseed, at the midpoint of the extraction cycle, refilling the cells in order to ensure better solvent contact with the oilseed. The cells containing the oilseed slide over rod-type screen plates with conical rods to minimize clogging of the screen plates. 14.5.6 New developments in solvent extractors Newer developments in solvent extraction plants are geared to very large extractors permitting 6000 to 10 000 Mt d–1 capacity on soybean and very efficient recovery of solvent.
14.6 Solvent recovery in oilseed processing 14.6.1 Recovery from miscella Hexane is the most frequently used solvent to extract oils. Other solvents have been considered (Hron et al., 1982; Johnson and Lusas, 1983). First step in recovering solvent from the miscella is to remove solid particles (fines). This can be done before or after solvent removal. However, if fines are left in the miscella and only removed after desolventization, they cause problems in evaporators and strippers, and fines remaining after desolventization contain more oil (that will be subsequently lost) than fines removed from the miscella. Solvent is usually removed from the miscella in a two-stage rising-film evaporator, followed by a ‘packed’ or a ‘disk and doughnut’ stripping column, which remove the last traces of solvent. The first-stage evaporator can be heated by hot solvent vapors from the marc desolventizer. These hot solvent vapors can also be used to preheat the miscella in a heat exchanger. The second-stage evaporator is heated by jacket steam. The solvent vapors from both evaporators are condensed whereas the oil from the evaporators, containing about 5% solvent, goes into the stripping column and is exposed to a counter-current flow of live steam to strip out the last traces of solvent. In the stripping column, the oil forms a thin film, giving it a large surface area as it flows over the packing, facilitating the stripping. Stripping columns are operated under high vacuum, 559–711 mm (22 to 28≤) Hg to help vaporize the solvent. Solvent and steam vapors from the stripper are condensed and pass to a ‘solvent/water separator’. The condensed liquids separate into two phases: hexane, being less dense, floating above the water. As the separator fills, excess hexane overflows from the top and goes to the solvent work tank. The separator height is chosen so that, at a predetermined liquid level, the weight of the liquid will force water at the bottom to pass up a siphon tube and go to a water stripper. There are still traces of hexane in the water from the solvent/water separator, so this water is scrubbed with live steam to ensure that no traces of hexane remain in water leaving the solvent plant. © Woodhead Publishing Limited, 2010
Separation technologies in oilseed processing 423 The steam from the water stripper is condensed and sent to the solvent/water separator tank. 14.6.2 Recovery from marc The marc is transported to a ‘desolventizer/toaster’, which is divided into several horizontal sections or ‘trays’ within a vertical cylindrical vessel. Sweep arms, attached to a vertical shaft skimming across the surface of each tray, keeps the marc agitated. The sweep arms also trip trap doors in each tray to allow marc to fall into the lower trays. Tray bottoms are steam jacketed and sometimes the vessel side walls also. Sparge steam, introduced into the top stack, helps to flash off the solvent. The solids absorb the condensed sparge steam, minimizing dust carryover to the condenser. The sparge steam also cooks or ‘toasts’ the solids. This helps to inactivate enzymes, like urease in soybean. Some desolventizers use a counter-current flow of live steam, introduced under the bottom tray and passing upward through perforations in each tray and through the marc on each tray. This greatly assists desolventization. The perforations are through hollow staybolts dispersed into each tray. The top tray has a steam-jacketed bottom. The heat emanating from the jacket helps flash off surface solvent in the incoming marc. Another design, a flash desolventizer, vaporizes solvent at low temperature to prevent any denaturing of the protein. The marc enters a recirculating stream of hot solvent vapor. Most of the solvent in the marc vaporizes. Some of the vapor, equal to what is flashing from the marc, is bled off through a rotary valve. These solvent vapors are then scrubbed to remove fines and liquefied in a condenser. The marc solids, still at about 1% solvent, are directed through a cyclone and rotary valve, then pass into a ‘flake stripper’ where the final traces of solvent are removed under low-heat, low-moisture conditions that will not denature or darken the protein. 14.6.3 Recovery from effluent air Air enters the solvent extraction system with the incoming material and through any leaks in the vessels, many of which are operating under a vacuum of about 2.4 to 5 mm (1 to 2≤) water column. This ensures that, if there are leaks, air will leak into the equipment rather than solvent vapors leaking out. This air will leave the solvent extraction system saturated with solvent vapor. Effluent air collected from all vessels likely to contain incoming air, such as the extraction vessel and the desolventizer, is directed into a common header, and then passes through a device designed to remove as much solvent as possible before the air is discharged. Two devices are preferred for maximum solvent recovery. One device is like a miniature solvent extraction system using mineral
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424 Separation, extraction and concentration processes oil as a solvent to absorb, and thus extract, hexane from the effluent air. The mineral oil is then stripped in a similar way to how vegetable oil is stripped. The stripping steam and hexane vapor are then condensed and pass to the main solvent/water separator. The hot mineral oil from the stripper is cooled and returns to the mineral oil scrubber. In this fashion, the mineral oil is continuously recirculated through the absorber, the stripper, and the cooler. The second device passes the vented air, still containing solvent vapor, through a bed of activated charcoal. Most of the solvent is absorbed by the charcoal while the air passes through. When the charcoal is almost saturated with solvent (and no longer effective), the incoming air is diverted to a second charcoal absorber, which is piped in parallel with the first one. The air from both absorbers is scrubbed with sparge steam to strip out the solvent. That solvent/water mixture is then condensed and sent to the main solvent/water separator. The absorber saturated with solvent undergoes a cleansing cycle to remove the absorbed solvent. Using either device, solvent loss from the processed oilseed through effluent air can be kept below 1.89 L t–1 (0.5 gal t–1), and, if the equipment is properly maintained and operated, solvent losses as low as 0.8–1.89 L Mt–1 (0.22–0.5 gal t–1) can be achieved. 14.6.4 Recovery from effluent water All mixtures of steam and solvent vapors, after being condensed, are sent to a common solvent/water separator where solvent overflows the top, and water siphons out the bottom, as described above. The water is then passed through a steam-heated ‘waste water stripper’ where live steam elevates the water temperature above the boiling point of the solvent but under the boiling point of water. The desolventized water then leaves the solvent extraction system.
14.7 Obtaining oil from fruit pulps Typical fruit pulp oils are olive oil and palm oil. These oils reside in the soft fruit pulp rather than in the kernel or seed. Palm fruit oil serves the same markets as other vegetable oils. Olive oil is valued for its flavor. 14.7.1 Palm oil Palm fruits grow in clusters on a central stalk. The fruit is composed of oily pulp within a tough outer skin and has seeds, or kernels, imbedded in the pulp. Two very different types of vegetable oil are obtained from this palm plant. Palm fruit oil comes from the pulp. Palm kernel oil comes from the seed. © Woodhead Publishing Limited, 2010
Separation technologies in oilseed processing 425 The clusters of freshly picked palm fruit are sterilized in steam chambers and then are sent through the stripping/threshing equipment to remove the individual fruits. The separated fruits are next washed to remove sand. Pressing is done in a twin-screw press to liberate the palm fruit oil. The twin screw press is operated at lower pressure than full-presses. The final step is to clarify the oil of moisture and debris. The twin-screw press is equipped with two parallel worm shafts that rotate in opposite directions and transport the fruit through drainage cages that surround the shafts. Hydraulically operated cone chokes apply back pressure. The liquid, a mixture of palm fruit oil and water, flows out through the drainage cage. Solid residue is pushed over the cone chokes by the action of the shafts. The expressed mixture of oil, water, and some solid impurities, is further processed to clarify the oil. The solids from the press consist of moist pulp solids, the outer skin, and the palm kernels. The solids are sent to a pneumatic separator to separate the kernels from the fiber and other debris. The kernels are cracked to separate the meats from the shells. The meats, still moist, are dried to 7% moisture and stored for subsequent full-pressing. 14.7.2 Olive oil Previous methods involved washing the olives, crushing them, stirring them into a thick paste, and then pressing to separate olive oil from the paste. Pressing was done in batch-operated plate presses by inserting bags between the plates followed by a hydraulic ram to squeeze out the oil. The oil was then centrifuged and clarified using diatomaceous earth in a subsequent filter press. Currently, centrifuges are used. The olives are first washed and then pass through a milling and beating stage followed by centrifuging. In the 1970s, three-phase centrifuges were used, employing hot water to obtain maximum oil yield. The wastewater stream presented ecological difficulties so, in the 1990s, the three-phase centrifuges were replaced by two-phase centrifuges, which greatly reduced the amount of wastewater (Harmsen and Mulder (2009, p. 20).
14.8 Future trends Recently, new objectives and new concerns have influenced oilseed processing, and equipment suppliers are developing new equipment and new procedures to address these concerns. Currently, most vegetable oils serving the food market are separated by solvent extraction because solvent extraction permits large capacity plants and maximum separation of oil. However, recent concerns about foods exposed to hexane and other chemicals, especially chemicals known to have deleterious consequences to the consumer, have © Woodhead Publishing Limited, 2010
426 Separation, extraction and concentration processes caused consumers to become suspicious of foods exposed to chemicals. Today, many foods are labeled ‘organic’ signifying that the foods have not been exposed to chemicals such as fertilizers, pesticides, solvents, and some preserving chemicals widely used in the past. This trend will probably continue, encouraging food producers to prepare foods without exposure to any chemicals. For that reason, future processors might favor screw pressing over solvent extraction (if screw pressing could be done at higher capacities and at lower cost) because no hexane or other chemical solvent is introduced into the food material being processed. Screw presses have traditionally been low-capacity machines compared with solvent extractors and consume more horsepower per tonne of oil extracted than solvent extractors. There is an incentive, today, for screw press manufacturers to develop higher capacity screw presses that consume less horsepower. Some manufacturers are already working on this. Other manufacturers are developing novel procedures, such as injecting liquid supercritical carbon dioxide through the oilseed within the screw press to maximize release of oil. This and other new developments have been discussed in this chapter. A new market is developing for the use of vegetable oil to replace petroleum products so that bio-renewable fuels (oils for diesel engines and ethanol from fermentation of the solid residue for gasoline engines) can diminish world dependence on fossil fuels. There has been some success producing ethanol from cereal grains such as maize and diesel oils from sunflower and rape seed. Success in these areas led to large-scale plants for making ethanol from maize starch and many small screw press plants to obtain vegetable oils for biodiesel. Some of the more efficient biodiesel plants also contain esterification equipment. Success in obtaining bio-renewable fuels from traditional food crops shows signs of placing greater demand on available food crops thereby raising the price for these food crops. This could have unfortunate long-term results. Also, the fossil fuel market is huge and could eventually consume a very large portion of available food crops. Therefore, efforts are currently being made to identify oil-containing materials that have not been processed before and could easily serve the bio-renewable fuel market without compromising the edible food market. A very promising source of untapped biorenewable oil is from algae. Algae are easy to grow, have sufficient oil content, and can, with some difficulty, be processed to obtain the oil. The world’s first algae biodiesel plant went online in Rio Hondo, Texas on April 1, 2008 (Cornell, 2008). It will make ethanol from the algae biomass first. Future plans involve also separating the oil. Algae has potential because, if and when proven successful, algae grown on only 9.5 million acres of land could replace all transportation fuels consumed annually in the USA. Farm cropping in the USA, in contrast, currently uses 450 million acres (Briggs, 2004). Recently, the New York Times reported that Exxon Mobil Oil Company, once skeptical of algae, has announced a 600 million US dollar program to produce biofuels from algae (Anon., 2009b).
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Separation technologies in oilseed processing 427
14.9 Sources of further information and advice Much has been written about standard equipment and procedures. There is no need to describe once again what is already well known and fully described in earlier books. For those interested in what was done in the past, The recovery of oils and fats from oilseeds and fatty materials (Shahidi, 2005) contains a thorough description of the early development of screw presses and solvent extractors and the procedures used from early development to state-of-the-art practiced at time of publication. Another good reference is Gunstone and Padley, (1997). Other good references are Modern technology of oils, fats & its derivatives, (NIIR, 2002) and Technologies for the recovery of residual oil (Harmsen and Mulder, 2009). Extrusion as preparation ahead of screw pressing and solvent extraction can also be further investigated through Farnsworth et al. (1986), and Lusas and Watkins (1988); Nelson et al. (1987), Williams (1993). Hexane is the most widely used solvent for solvent extraction of vegetable oils. Information about other solvents that are used can be found in Shahidi (2005) and Hron et al. (1982). Some manufacturers of screw presses, solvent extractors, and preparation equipment are identified in this chapter. Others can be found in Anon. (2009a) and on the Internet. Suppliers of plant and equipment Anderson International Corp, 6200 Harvard Avenue, Cleveland, OH 44105, USA (http://www.andersonintl.net). Crown Iron Works Company, P.O. Box 1364, 1600 Broadway Street N.E., Minneapolis, Minnesota 55440-1364, USA (http://www.crowniron. com). Desmet Ballestra Group N.V. Minervastraat 1, B – 1930 Zavetem, Belgium (http://www.desmetballestra.com). The Dupps Company, 548 N. Cherry Street, Germantown, OH 45327, USA (http://www.dupps.com). Duyvis, B. V. Machinefabriek P. M. Duyvis, 1541 KD Koog ann de Zaan, P.O. Box 10, The Netherlands (http://www.pmduyvis.nl). The French Oil Mill Machinery Co., 1035 West Green Street, P.O. Box 920, Piqua, Ohio 45356-0920, USA (http://www.frenchoil.com). Harburg Krupp Maschinentechnik GmbH, Werk Harburg, Postfach 900880, Seevestrasse 1, D-2100 Hamburg 90, Germany (http://www.harburgfreudenberger.com). Insta-Pro International, 10104 Douglas Avenue, Des Moines, IA 50322, USA (http://www.insta-pro.com). Lurgi GmbH, Lurgiallee 5, ED-60259 Frankfurt am Main, Germany (www. lurgi.com).
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14.10 References Anon. (2009a) NY Times Available from: http://www.nytimes.com/2009/07/14/business/ energy-environment/14fuel.html (Accessed 15 July 2009). Anon. (2009b), ‘Plant and equipment survey’, Oils & Fats International, 25, 38–41. Baer S, Williams M and Zies C (1966), ‘Pre-treatment of oleaginous plant materials’, US Patent 3,255,220. Briggs M (2004), Wide scale biodiesel production from algae, University of New Hampshire, Physics Dept. Available from: http://www.unh.edu/p2/biodiesel/article_alge.html (revised August 2004) [Accessed 16 July 2009, note algae misspelled] Cornell C (2008), First algae biodiesel plant goes online: April 1, 2008, Available from: http://gas2.org/2008/03/29/first-algae-biodiesel-plant-goes-online-april-1-2008/ (Accessed 16 July 2009). Farnsworth J, Johnson L, Wagner J, Watkins L and Lusas E (1986), ‘Enhancing direct solvent extraction of oilseeds by extrusion preparation’, Oil Mill Gaz, 91, 30. Friedrich J, List G and Heakin A (1982), ‘Petroleum-free Extraction of Oil from Soybeans with Supercritical CO2’, J Am Oil Chem Soc, 59, 288. Friedrich J and Pryde E (1984), ‘Supercritical CO2 Extraction of Lipid-bearing Materials and Characteristics of the Products’, J Am Oil Chem Soc, 61, 223. Gunstone F and Padley F (1997), ‘Extraction of lipids from natural sources’, Lipid Technologies and applications, New York, Marcel Dekker, Inc., 113–135. Harmsen P and Mulder W (2009), Technologies for the recovery of residual oil, Available from: http://www.york.ac.uk/res/sustoil/Pages/Deliverable%202.3%20FINA.pdf (Accessed 16 July 2009). Hron R, Koltun S and Graci A (1982), ‘Biorenewable solvents for vegetable oil extraction’, J Am Oil Chem. Soc 59, 674A. Johnson L and Lusas E (1983), ‘Comparison of alternative solvents for oils extraction’, J Am Oil Chem Soc 60, 229. Lusas E and Watkins L (1988), ‘Oilseeds: Extrusion for Solvent Extraction’, J. Am. Oil Chemists’ Soc. 65, 1109. Nelson A, Wijeratne I, Yeh W, Wei T and Wei L (1987), ‘Dry extrusion as an aid to mechanical expelling of oil from soybeans’, J. Am. Oil Chemists’ Soc. 64, 1341. NIIR (2002), Modern technology of oils, fats & its derivatives, New Delhi, National Institute of Industrial Research, India, ISBN: 8178330857. Radio Iowa, ‘Elsworth company breaks ground for soy plant,’ Available from: http://www. radioiowa.com/gestalt/go.cfm?objectid=57F7E7F2-B31D-4124-ADC4E78B4C8934B1 (Accessed 16 July 2009). Shahidi F (2005), ‘The recovery of oils and fats from oilseeds and fatty materials’, Bailey’s industrial oil and fat products, Sixth Edition. Edited by Fereidoon Shahidi, New York, John Wiley & Sons, Vol. 6, 2589–2678. Vadke V, Solulski F and Shook C (1988), ‘Mathematical simulation of an oilseed press’, JAOCS, 65, 1610. Also available from: http://www.springerlink.com/content/ b71x7j050272mvjn/ (Accessed 16 July 2009). Voges S, Eggers R and Pietsch A (2008), ‘Gas assisted oilseed pressing’, Sep Purif Technol, 63, 1–14. Willems P (2007), Gas assisted mechanical expression of oilseeds, Universiteit Twente, Nederland. Available from: http://doc.utwente.nl/58041/(Accessed 16 July 2009). Williams M and Baer S (1965), ‘The expansion and extraction of rice bran’, J. Am. Oil Chemists’ Soc. 42, 151. Williams M (1983), ‘Description of Anderson International’s vector shaft analysis technique’, Oil Mill Gaz, March, 36–37.
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Separation technologies in oilseed processing 429 Williams M (1990), ‘Apparatus and method for the continuous extrusion and partial deliquefaction of oleaginous materials’, to Anderson International Corp. US Patent 4,901,635. Williams M (1991), ‘Extruded starter pig feeds’, Feed Manage 42, 20. Williams M (1993), ‘Preparation of oilseeds to improve extraction of fats’, Extrusion Commun 6, 12.
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15 Separation technologies in brewing G. J. Freeman, Campden BRI, UK
Abstract: The series of separation processes that characterises the brewing process is overviewed with emphasis on those processes that provide opportunities for technological progress. Extraction of raw materials in the brewhouse and the filtration and inertial separation processes employed are described. Yeast separation after fermentation is mainly achieved through natural flocculation processes, which may be enhanced by fining agents. Filtration to commercial quality beer clarity may be achieved by filter aid filtration, including novel filter aids that are more environmentally friendly, or membrane filtration. Other applications of membranes include control of dissolved gas levels and recovery and reuse of detergents. Key words: wort separation, yeast flocculation, beer filtration, membrane technology.
15.1 Introduction The brewing process is characterised by a sequence of many separation processes. The processes of brewing (hot extraction), fermentation, maturation and end processing generally comprise two or three phases of solid, liquid or gas. Processing requires accurate control of the appropriate variables in order to achieve the desired quality and consistency in final product. Sales of beer globally have become dominated by stable products with long shelf life, perhaps as long as a year. This requires a high standard of hygiene, process equipment and operating procedures. Shelf life is limited by either microbiological, flavour or colloidal (turbidity) instability. Brewers are fortunate in that no pathogens can grow in mainstream beer products. However, there are a variety of micro-organisms that can spoil beer by causing
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Separation technologies in brewing 431 turbidity, acidity and other off-flavours (O’Rourke, 2000). Flavour instability is a complex deterioration of the product by a variety of chemical reaction schemes. ‘Stale’ beer is characterised by the loss of desirable bitterness and a variety of undesirable flavours such as ‘wet cardboard’ and ‘winey’ (Bamforth, 2000). Colloidal instability may occur in a variety of ways, but the most important, and that which practically all beer is susceptible to, is the production of particles by flocculative interactions between polyphenols and certain polypeptides (Siebert, 1997). The microbiological, flavour and colloidal stability is extremely adversely affected by the presence of oxygen in the final product. There is a need to exclude oxygen as much as possible after the process of primary fermentation. Final packaged beer should if possible have as little as less than 0.1ppm of oxygen. Again, this necessitates rigorous process plant design and procedures. Most beer is pasteurised, and this, along with hygienic plant design and operation, normally achieves microbiological stability over the shelf life of the product. Flavour stability may perhaps be quantified by the measurement of concentrations of ‘marker’ chemical compounds (Malfliet et al., 2008) but this is complex and still under development. Therefore, colloidal stability is usually employed as the definitive measure for product shelf life. There is a need to process the polyphenols and polypeptides that are responsible for colloidal instability by allowing the formation of the particles and then removing them by a variety of separation process options. Indeed, the more effective the separations throughout the process, even very early in the process, the better the colloidal stability of the beer.
15.2 Characteristics of brewery products There are a large number of distinctive beer styles. They are characterised not just by region of origin, but by differences in process and packaging style that are particularly relevant in this chapter. Most beer in the world is retailed in small packs of glass bottles, stainless-steel or aluminium cans or plastic (most commonly polyethylene terephthalate) bottles. These products are normally filtered to an appealing visual clarity. Cans and glass bottles are most commonly pasteurised in package (tunnel pasteurisation). This is not always possible with plastic bottles, although in recent years some bottles amenable to high-temperature processing have come on to the market (Martin, 2002). Plastic bottles are most commonly sterile filled after sterile filtration. Larger containers usually employed in bars and restaurants are varieties of keg. Microbiological stabilisation is not possible in these containers, but is achieved by sterile filtration or, more commonly, by in-process pasteurisation (flash pasteurisation by plate heat exchanger). In-process pasteurisation is more thermally efficient than in-pack. Thus, there are drivers to employ flash pasteurisation in these times of increasing energy prices (Browne, 2008). © Woodhead Publishing Limited, 2010
432 Separation, extraction and concentration processes Some beers, normally originating from discrete regions, are characterised by not having been filtered. Cask-conditioned beers undergo secondary fermentation (maturation) in the container. The yeast is sedimented from the beer by employing flocculents known as finings. Bottle-conditioned beers undergo secondary fermentation in the bottle.
15.3 Selection of technology and raw materials appropriate to brewery products The main raw material employed in beer production is malted barley. The malting process involves initiating the growth of the barley plant induced by increasing the content of water by steeping. The grain is allowed to grow (germination) for a period of some days, which causes structural and enzymic changes in the grain that enable the malt to be processed in the brewhouse. These changes, known as modification, are allowed to occur up to a certain point dependent on the specification required for the beer product. Finally, the process is stopped by drying with hot air (kilning) that produces a stable product. The time–temperature cycle employed in kilning also controls aspects of the malt character notably flavour and colour. Malts that are employed to make darker beers such as ales are typically more modified and kilned to higher temperatures. In the brewery, the malt is milled with either roller mills or hammer mills. The detailed choice of mill depends upon the extent of modification of the malt and the technology employed in the brewhouse. The milled malt, known as grist, is contacted with warm water in a process known as mashing. The grist may be supplemented with other processed cereals such as wheat or maize. In the mash, several biochemical reactions occur as a result of the enzymes in the grist. The result is a sugar- and nutrient-rich medium known as wort. The wort is separated from the remaining undissolved brewers’ grains by filtration through the bed of grains itself. The technology selection is limited by the choice of raw materials as described herein. The wort must undergo a boiling process. This is required for reasons of flavour, production of bitterness and flavour from hops that are traditionally added to the boil and for stability and sterilisation reasons. Sugar-rich syrups may also be added to supplement the wort. Excess protein precipitates as large loose solids known as trub that requires removal along with the remaining hop solids. This separation process is most commonly performed in a vessel known as a whirlpool described herein. The wort is then cooled, transferred into the fermenter vessel and pitched with active yeast. The fermentation process is exothermic and is controlled by a temperature programme. Some darker beers such as ales and stouts may be allowed to reach temperatures greater than 20 °C, whereas typical lagers are not fermented above 15 °C. Thus, the lager fermentation takes longer
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Separation technologies in brewing 433 but produces a more delicate flavour, an effect that is exaggerated by the employment of less kilned, less flavoursome malt products. Separation of most of the yeast occurs through the natural flocculation and either sedimentation or flotation of the resultant flocs. A secondary fermentation is required for reasons of flavour. Most beer then undergoes a cold stabilisation (storage) period. This is required to precipitate more excess protein by flocculation with polymerised polyphenols. This is required for products that are served chilled (most beer) and/or require extended shelf-life. Most beer is clarified by fine filtration. Some employ flocculents in package, known as finings, that maintain suspended solids in a large floc away from the dispense point. Cask-conditioned ales are the main example of this. Bottle-conditioned beers are similar, also containing live yeast as a sediment but without finings. Microbiological stabilisation is required if yeast is not present in package to inhibit micro-organism growth. This occurs by pasteurisation or sterile filtration before packaging.
15.4 Wort production in the brewhouse After the biochemical reactions have occurred in the mashing process, the liquid wort is separated from the brewer’s grains by filtration through the grain bed itself. The majority of the extract is recovered to the wort by rinsing (sparging). The filtration rate, which commonly limits the capacity of the brewhouse, is mainly controlled by both the coarseness of grind and the depth of the bed. The three dominant wort separation technologies are mash tuns, lauter tuns and mash filters. The relative merits of the three processes have been reviewed by O’Rourke (1999) and are described hereafter only in broad classification terms. A mash tun is characterised by mashing and wort separation occurring in the same vessel. The grain and hot water are mixed in such a way that air is entrained and the grain bed floats. Sparging occurs by spraying hot water on top of the bed. The mash tun is limited to highlymodified, coarsely ground malt (a fraction may be from other cereal starch sources) that makes it most suitable for the production of ales. The coarse grind gives a thick mash and a deep filter bed (1 m deep). A lauter tun involves similar technology to a mash tun, but the mashing process proceeds in a separate vessel called a mash mixer before pumping into the lauter tun. The lauter tun is more flexible on the malt qualities that may be employed and, for example, is more appropriate to many lagers that are made from less well modified malts. The grind is less coarse than for a mash tun and hence the bed must be shallower (40 cm). Flexibility in choice of raw materials is further improved by the employment of raking of the bed if the filtration rate is found to be too slow. Mash filter technology has been greatly advanced over the last three decades. The grain is ground to a relatively fine flour by a hammer mill (most commonly) as opposed to roller mills employed for © Woodhead Publishing Limited, 2010
434 Separation, extraction and concentration processes mash tuns and lauter tuns. The fine grind necessitates the need for a thin filter bed. The advances have mostly centred around pressurisation of the filter bed to increase filtration rate by, for example, pneumatically driven polymeric sheets (Eyben et al., 1989). The lauter tun and mash filter technologies allow much greater throughput than mash tuns. Mash filters have been proposed as part of a continous brewing technology, being operated semi-continuously with one in forward flow with the other being readied for use (Harmegnies et al., 2003).
15.5 Whirlpools and applications in brewing There is a need to remove solid material after boiling the wort in the brewhouse. Excess protein is removed as precipitated trub, which if it remains through to fermentation will cause sulfurous off-flavours and complicate downstream processing. Most breweries employ whirlpools, the mechanism of which is described in Fig. 15.1. The vessel is filled through a tangential main and then allowed to stand, the liquid maintaining a rotational flow, for a period typically of 15–60 min. Clear wort may then be drawn off away from the accumulated solids at the bottom centre of the vessel. The trub solids exist as loose, delicate flocs. With all flocculation processes, it is necessary to ensure that the correct shear forces are applied in the transferring pipe mains and pumps. Excess mixing results in small flocs that are difficult to separate, typically manifesting as a poor ‘cone’ of
Continuing liquid rotation after vessel is full causes concave liquid surface
Resultant liquid flow. Upflow at centre insufficient to lift accumulated solids
Tangential inlet(s) for fill.
Force due to larger liquid head nearer side of vessel
Force due to particle density being higher than wort density
Force balance causes solids accumulation bottom centre
Fig. 15.1 Mode of operation of a whirlpool for trub separation.
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Separation technologies in brewing 435 accumulated solids smeared across the base of the vessel. Inadequate mixing means that the trub particles do not contact frequently enough, resulting in similar separation problems (Virden, 1995).
15.6 Yeast flocculation and applications in brewing The traditional brewing process relies on the timely flocculation of yeast in order to clarify the beer sufficiently so that downstream processes are able to finish the beer to consumer satisfaction. Thus, it has always been the case that a major selection criterion for a suitable brewing yeast is that it flocculates at the correct stage of the fermentation (Soares, 2009). The application of modern centrifuges means that latterly it has been possible to employ less flocculent yeast strains. The flocculation process itself comprises interaction between polysaccharides and glycoproteins on the cell surface (Evans and Kaur, 2009). The timing of rapid flocculation is controlled by factors such as the presence of calcium ions, a low level of unfermented extract and temperature. Some relatively hydrophobic strains, such as those that are traditionally employed to make ales, may be induced to flocculate to the surface of the wort to be removed by ‘skimming’. Most commonly, brewing yeast strains flocculate to the bottom of the vessel. It is in the nature of all flocculation processes that they are affected by mixing in the system. The flocculent particles require a degree of mixing so that they contact and join together to form flocs. Mixing is provided in the fermentation vessel by the evolution of carbon dioxide. However, it has been suggested that a degree of forced mixing is beneficial (Boulton, 2009). Pumping the fermenting wort around a loop around the vessel appears to result in: more efficient cropping of the yeast (presumably making downstream processing less expensive also), reduced fermentation cycle times, increased ethanol yield and the ability to cool the fermentation more rapidly at the end of the cycle because of improved heat transfer. One interesting phenomenon that is currently under intense investigation is premature yeast flocculation (PYF). This effect has probably always intermittently appeared without being correctly diagnosed. The consequences are serious, with the yeast unable to complete fermentation resulting in excessive residual carbohydrate and other undesirable beer flavours (Lake and Speers, 2008). At the time of writing it is still not well understood. Most likely, the problem originates from microbial, probably fungal, contamination of barley in the field (Evans and Kaur, 2009). The PYF-inducing factor, or possibly more than one factor, appears to have molecular weight less than 100 kDa. However, the effect of process factors in both malting and brewing are unclear. In the germination process during malting, air temperature, air flowrate and humidity have been implicated as conditions that affect the presence of the PYF factor. However, malts that induce PYF in some brewing
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436 Separation, extraction and concentration processes processes may not in others. There are probably differences caused by the final wort quality and the conditions in the fermentation vessel.
15.7 Beer fining agents Finings are flocculents in the brewing process. The most important is isinglass. This is a remarkable substance that is manufactured by solubilising the swim bladders of fish into dilute mineral acid. The result is a viscous suspension of almost pure collagen. When added to beer, the collagen molecules adopt a net positive charge. The large majority of particles in the beer possess an overall negative charge. This is particularly true of yeast cells that possess a very significant negative charge as long as they are still alive. Hence, the collagen molecules act as a coagulant, neutralising the beer particulates’ surface charge and allowing them to overcome repulsive forces. The collagen molecules are long macromolecules with a triple-helix structure that are able to bridge between the beer particulates. The consequence of these interactions is that the beer particulates form large flocs that are able to sediment quickly. Isinglass is employed usefully in both beer that is matured in the brewery and in beer that is matured in cask. It may be added to beer in the cold storage vessel. This has the effect of speeding up the process, reducing costs and increasing brewery capacity. There is also a superior partition of the solid and liquid phases. The bulk of the beer (the ‘supernatant’) will be faster and less expensive to process downstream, notably in the filtration operations. Also, flocculents are capable of removing very small particles that even filtration processes may allow to pass. Finings therefore have a beneficial effect on colloidal stability. As described earlier, isinglass finings are essential for the clarification of unfiltered cask-conditioned beer. The cask undergoes a suitable period of secondary fermentation. After placement in the retail outlet, the finings will have clarified the beer after typically settlement for 24 h. The finings suspension is viscous compared with the beer and also comprises only a very small fraction of the volume of the beer. Thus, there is a challenge to achieving a homogeneous mix to optimise performance. The finings are often added in-line during tank transfer, perhaps on a pipe bend to maximise turbulence or the turbulence of fill is used in cask. However, one study has identified an optimised mixing regime comprising a short, relatively intense mix followed by a longer, more gentle mix, ideally achieved by two static mixers (Freeman et al., 2003). The performance of isinglass finings may be enhanced by materials called auxiliary finings. These are polysilicates, polysaccharides or a mixture of both. Addition to beer is before the addition of isinglass, and best performance requires good mixing of the beer and auxiliary finings. Auxiliary finings form a negative charge at beer pH. Hence, they support the performance of © Woodhead Publishing Limited, 2010
Separation technologies in brewing 437 the isinglass by ensuring that the beer particulates adopt a negative charge. This is especially true for yeast cells; if the cells are dead, and typically a moderate percentage are, they lose their surface charge. An alternative to isinglass has been identified based on pectin (Duan et al., 2008). It is claimed to perform as well as isinglass, and would have the commercial advantage that the resultant beer would remain indisputably suitable for consumption by vegetarians.
15.8 Filter aid filtration and applications in brewing Filtration in breweries is most commonly accomplished by the use of filter aids. These substances, used as slurried powders, form incompressible and highly porous filter beds, thus allowing the relatively free flow of beer. In small plants with small batch sizes, simpler filtration technologies may suffice. The most common filter aid used in breweries is kieselguhr or diatomaceous earth (Fig. 15.2). These materials comprise fossils or skeletons of microscopic saltwater or freshwater life known as diatoms. When they die they sink and form deposits that are mined, processed and size-classified to give kieselguhr of various grades. The disadvantages of kieselguhr are that it is a health hazard (by dust inhalation) in its dry form as delivered to the brewery and that it is in itself non-biodegradable, with a concentration of organic solids, and is thus expensive to dispose of in landfill sites. The configuration of a filter aid filtration system is shown in Fig. 15.3. Before processing occurs a precoat of filter aid is deposited onto the filtration
Fig. 15.2 An electron micrograph of kieselguhr (diatomaceous earth).
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438 Separation, extraction and concentration processes Buffer tank
Filter aid slurry tank
Base of cold storage tank Dosing pump
‘Rough’ beer
Filter
Pump
Filtered beer
Recycle and precoat facility
Fig. 15.3 A diagrammatical representation of a filter aid filtration system.
surface. This is achieved by recycling of a water/filter aid slurry around the filter. After several minutes, the precoat is deposited completely onto the filtration surface and the recycling water is clean. The precoat is necessary to ensure efficient filtration of the early part of the beer run, to guarantee the integrity of the filter throughout the run and to aid removal of the filter cake after the process cycle. After precoating, the filter is smoothly put into ‘forward flow’ mode. The filter aid slurry is added continuously to the flowing beer stream. Thus, the filtration surface is constantly being regenerated. In this way, the filtration run time is extended causing the process to be commercially viable. With regard to the actual filter unit, options may be divided into plate and frame type, leaf type and candle type. Plate and frame filters have been a workhorse of the brewing industry since the inception of filter aid filtration. They are known to enable excellent filtrate clarity. However, they are not amenable to full automation causing long down time between filter runs and an increased manpower requirement. Most brewers would now choose leaf or candle filters for which beer recovery, cleaning and re-starting may be automated by process control systems. Selection of leaf or candle is a brewery specific decision. Leaf filters are mechanically more complex and higher maintenance, but have more flexibility in flowrate and are not as vulnerable to process interruption. The bulk filtration duty in a brewery is a demanding unit operation. It is essential for product clarity, and also for colloidal stability. It should significantly lower the quantity of contaminant micro-organisms presented to the pasteuriser, because heat should be used sparingly if flavour impairment is to be avoided. If sterile filtration is employed, the bulk filtration stage must still give a high degree of clarity because the majority of sterile filtration systems have very limited solids-holding capacities. In this chapter, alternative filtration technologies and filter aids are
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Separation technologies in brewing 439 considered. These alternatives are challenged by the intrinsic difficulties of beer as a product for filtration. The low temperature (0 °C) and presence of dissolved solids and alcohol means that viscosity is quite high (at least 2 mPa s). Of even more significance is the nature of the suspended solids. These may be present in very high levels, perhaps up to 0.2% by volume or even higher over short periods during tank run-off (Fig. 15.4). Practically all of the suspended beer solids are compressible, which causes them to form filter cakes impermeable to beer flow. Many are small, meaning that filtration must be through fine, small flow channel media to achieve acceptable clarity. Filtration may also be impaired by colloidal substances such as b-glucan gels (Waiblinger, 2002). Alternative filter aids (and alternative technologies) are currently in use. Perlites consist of thermally expanded volcanic glass (Davies, 2004), crushed to form microscopic flat particles (Fig. 15.5). Perlites are less efficient filter aids than are kieselguhrs but are perceived as being safer than kieselguhr. However, it lacks the remarkable skeletal structures of the diatoms that comprise kieselguhr. As a consequence its filtration performance is not as good. In order to achieve the required filtration performance, secondary filtration (e.g. sheet or cartridge filters) is required. It is interesting to compare the filtration performance of kieselguhr and perlite in some detail (Fig. 15.6). This graph shows the particle sizes remaining in beer after filtration with kieselguhr and perlite (of similar permeability). Note that the kieselguhr exhibits a very exact particle size at which almost all smaller particles will pass through and almost all larger particles will be retained. This is not the case with perlite, which although it removes more very small particles than Yeast cells
25 20 15 10 Protein-polyphenol particles 5 0 0.5 0.6 0.6 0.7 0.8 0.9 0.9 1.0 1.1 1.3 1.4 1.5 1.7 1.8 2.0 2.2 2.5 2.7 3.0 3.3 3.6 4.0 4.4 4.9 5.4 5.9 6.5 7.2 7.9 8.7 9.6 10.6 11.7 12.9 14.2 15.6 17.2
Volume in size class (mm3 ml–1) ¥ 10–6
30
Particle diameter (microns)
Fig. 15.4 The particle-size and concentration distribution in a sample of pre-filter beer.
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440 Separation, extraction and concentration processes
Fig. 15.5 An electron micrograph of perlite.
Volume of particles in class (ml l–1)
4.5 4.0 3.5 3.0 2.5 2.0 1.5 1.0 0.5
16.0
10.1
(mm)
6.4
Particle d iameter
4.0
2.5
1.6
1.0
0.6
0.4
0.0
Unfiltered beer Perlite filtered Kieselguhr filtered
Fig. 15.6 Particle-size distributions and concentrations of samples of beer before filtration and after filtration with perlite and kieselguhr.
kieselguhr, it allows particles in excess of 1.5 mm to pass into filtered beer. The two filter aids operate in a very different manner. The kieselguhr behaves somewhat like a sieve. The perlite, however, filters more like a depth filter (mass filter). The performance of perlite shown in Fig. 15.6 is less desirable than the kieselguhr filtration. The larger particles will manifest themselves as a more obvious, visible haze in the beer. Also, a higher proportion of any
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Separation technologies in brewing 441 spoilage bacteria will pass through the perlite filter. Hence, there is a need for secondary filtration or an increase in pasteurisation intensity. The latter may impinge adversely on beer quality. It is likely that many of the novel, alternative filter aids now on the market will not filter beer to the same standard of clarity that is possible with kieselguhr. This is because of the unique structures in kieselguhr that cannot economically be reproduced artificially. The advantages of the alternative filter aids must be one or more of: a stabilisation effect, the ability to be regenerated and reused, and improved health and safety or environmental factors. After the main filtration operation, many breweries employ secondary filtration operations. Often the purpose is to provide sparkling clarity and enhanced shelf life before the beer becomes turbid in-pack. Some products are sterile filtered. This means that essentially all microbes with the potential to spoil beer are removed. The most common technologies installed for these purposes today are sheet filters (Brunner, 1987) and cartridge filters (Tubbs, 1998). Sheet filters mainly consist of cellulosic fibres that have been compressed into a thin mat and arranged in a plate and frame filter press. Cartridge filters comprise small units enclosing a filter element commonly of polymeric fibre sheets that are pleated. The demands on beer quality today, especially if sterile filtration is required, mean that, on occasions, three or four filtration steps are performed.
15.9 Regenerable and reusable filter aids and applications in brewing It is clear that kieselguhr would become much more environmentally acceptable if it were to be reused in the process. Alternatively, the spent filter cake could become a useful co-product of the brewing process. It has been suggested that spent kieselguhr could be employed for enhancing the nutrient value and structure of agricultural soil either directly or by composting first (Russ, 1993). However, in recent years, this use has become unfashionable because of the perceived health risk. It is possible to add spent kieselguhr to construction materials such as bricks and tiles. The problem here is economic in that individual breweries do not produce enough to justify the transportation to central processing. Hence, there is still scope for engineers to develop processes that enable kieselguhr to be reused in the brewery. There are established technologies involving sodium hydroxide and also furnacing (Russ, 1993). In both cases, economy of scale suggests the need for removal of spent kieselguhr from the brewery to a centralised processing station. Recently, studies have been performed on the use of hydrocyclones to separate the filter aid particles from organic material (Poku, 2004). Hydrocyclones operate by the conversion
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442 Separation, extraction and concentration processes of pressure energy into vortex flow, thus enabling separation of relatively dense particles from liquid suspension and less dense particles. Filter aid particles may be concentrated by a factor of up to twenty-five (depending on particle size) in one pass. Yeast cells are not concentrated and therefore are effectively washed out from the spent filter aid. There is the possibility of utilising several in series to effect a good separation, but there will still be a need for chemical washing of some stubborn organic residues. With all the available technologies the skeletal structure of kieselguhr means that only partial recycle is possible. There will be loss of material owing to particle attrition. There have been a variety of synthetic, regenerable filter aid filtration systems proposed. One system is based on a synthetic polymer (Brocheton et al., 1995). It is granular with a typical particle size of 35 mm, which is larger than the kieselguhr grades normally employed. However, the particles are claimed to be hydrodynamically ‘lighter’ than kieselguhr, which assists in the development of smooth, even filtration cakes. Regeneration is by hot caustic solutions, this means that the filter aid may be blended with polyvinyl polypyrrolidone (PVPP) and regenerated together. PVPP is a polyamide that adsorbs the polyphenolic component in beer that is responsible for combining with certain polypeptide fractions and polymerising into visible particles. Thus, it extends beer shelf-life. Stabilisation (of polyphenolic sources of instability) and filtration are achieved in one unit operation. Another system employs a mixture of synthetic microballs (for filtrate clarity) and fibres (for cake flexibility) (Harmegnies et al., 1997). Similarly to the previous system, the main regeneration process is with a hot caustic solution and therefore PVPP may be incorporated into the filter aid to provide stabilisation and filtration. A major chemical company has recently commercialised a regenerable filter aid comprising 30% by mass PVPP on a matrix of polystyrene (Zimmermann et al., 2008). In order to maintain the permeability of the filter bed, the particle size of the particles is larger than that of an equivalent kieselguhr. Conceptually, this would result in poorer filtered beer clarity, but the fact that the material is regenerable justifies a deep precoat layer and higher bodyfeed dosage rate enables acceptable beer clarity. Employment of candle filtration technology has been recommended with this filter aid (Ferstl and Zuber, 2009). With all three systems, filtrate quality and an evaluation of economic worth were both positive. Kieselguhr suffers from the disadvantage of transmission of oxidising transition metal ions into beer (notably iron) that the regenerable filter aids do not. Disadvantages of regenerable filter aid systems are the need for precoating and bodyfeed mixtures to be identical and also there is very limited flexibility over the dosage rate of stabiliser if it is employed. The economic benefits would depend on the current process. The required process modifications may be minor or substantial, thus affecting the economic attractiveness of the process. Pilot trials are recommended to assess process variables such as dosage rate (and consequently run length) and frequency of regeneration.
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Separation technologies in brewing 443
15.10 Bulk beer filtration by membranes In recent years, an acceptable alternative to filter-aid filtration has been developed that should reduce costs and improve the environmental image of brewers. This technology is cross-flow (or tangential flow) microfiltration by membranes. The challenge facing membrane filtration is the presence of particles in the micrometre and sub-micrometre ranges (Fig. 15.4) that cause severe membrane fouling. A partial solution is to pump the unfiltered beer (known as retentate in cross-flow processes) across the membrane surface, inducing a scouring effect. However, in practice, this is only partially successful because of the presence of small particles, colloids and macromolecules (Fig. 15.7). Nevertheless, the technology has been developed to the extent that the process economics are now viable. Currently, there are three commercially available systems for primary filtration of large volumes of beer. Norit supply a tubular membrane system (Noordman et al., 1999). The cost of replacement membranes is an important operating cost, and the Norit membranes have been shown to last for over 500 filtrations. The flow rate performance of the membranes is improved by several ‘flux-enhancement’ techniques. As the flowrate through a module falls, the pressure drop across the membrane is gradually increased, thus increasing the fouling rate and shortening the filter run, but maintaining an acceptable flow rate. Another technique is to employ ‘back flushing’. This entails increasing the pressure on the filtrate (known as permeate in crossflow processing) so that it exceeds the pressure of the retentate for a few seconds every, say, two hours. Although there is therefore a loss of acquired permeate, the effect of the backflush is to remove some of the fouling layer back into the retentate. This temporarily improves the flow rate of permeate after backflush. Norit have also developed cleaning regimes that include enzymes as well as detergents.
Molecular gel layer Yeast
In depth pore plugging Macromolecules
Cake formation
Surface adsorption Fine particles
Fig. 15.7 Diagrammatical representation of the membrane fouling mechanisms in crossflow microfiltration of beer.
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444 Separation, extraction and concentration processes The ‘Profi’ system is supplied by a collaboration between Westfalia (part of GEA) and Pall Seitz Schenk (Anon., 2004). Again, the membrane format is tubular. However, this system is characterised by a high-performance disc-stack Westfalia centrifuge immediately upstream of the membranes. The centrifuge increases the membrane flow rate, but, because the fouling is mostly caused by small particles that are inefficiently removed by centrifugation, the performance benefit is actually quite small (0–50%). The main benefit is reduction of the final volume of retentate before the suspended solids concentration is too high to allow efficient processing. Essentially, beer recovery from slurry operations has been moved upstream of filtration. Indeed, it is claimed that, in at least some installations, there is little or no recycling of retentate (Anon., 2009). The membranes are regenerable with caustic and oxidising detergents, and are sanitisable up to 80 °C. A third system is supplied by another collaboration between Alfa Laval and Sartorius (Borremans and Modrok, 2003). Again, the membranes are preceded by a high-performance (Alfa Laval) centrifuge. However, the Sartorius membranes are in a very different format to the systems described above. They are supplied as cartridges of square membrane flat sheets. They are separated on the retentate side by turbulence promoting gauze, which helps the cross-flow effect but necessitates the use of a centrifuge to prevent blocking of the cartridges on the retentate side. Cleaning is claimed to only require rising with water and caustic solution. The membranes comprise a porous structure that is asymmetric. For cross-flow processes, this is optimal when the narrow pore diameter is on the retentate side. The benefit of using the centrifuge is debatable. With the latter two systems, the power supplied to operate the centrifuge is at least partly recovered by a lesser need for a high cross-flow velocity over the membranes and a consequent reduced refrigeration demand. There will be a small increase in temperature by centrifuge processing, perhaps 1 °C, which is potentially slightly damaging to the colloidal stability of the beer owing to re-solution of some of the beer haze particles. It is worthy of note, however, that centrifuges have improved significantly in recent years, with improved particle removal performance and increasingly sophisticated and reliable sealing systems (Meckler, 2003). One of the main developments that has enabled improved performance is the employment of polyether sulfone (PES) membranes. These were until recently difficult to manufacture with the required microporous structure and with sufficient repeatability, but owing to an important development this is no longer the case (Riddell, 2002). For the filtration of beer they exhibit excellent low-fouling properties. The disadvantage of polymeric membranes compared with ceramic membranes, for example, is that replacement is required more frequently. The polymeric membranes last probably at least 500 hundred process runs, perhaps equivalent to two years operation. All three of the main commercially available systems employ PES membranes. The cross-flow aspect of the technology makes the plant more viable on a
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Separation technologies in brewing 445 large scale. On a small scale, membrane filtration is commonly employed for sterile or polishing filtration, often with a cartridge filter. The membranes are supplied in modular form. This has the benefit that membrane cartridges may easily be replaced when required without interruption to process. Also, some modules may be undergoing cleaning while others are processing beer. This raises the possibility of continuous processing. Cross-flow microfiltration is more amenable to automation than filter aid filtration and so there is the potential for a continuous process that requires minimal process supervision. However, a recent installation has shown more need for supervision than was anticipated (Pickerell and Heeb, 2008). Operator experience has been required in the operation of the centrifuge, control system and instrumentation maintenance, although this is an area for improvement and development in the future. Beer stability is very sensitive to the presence of oxygen as a source of free radicals that cause the beer to taste stale (oxidised) and have a deleterious effect on colloidal stability and thus shelf-life. A major cause of oxidative damage are the ions of transition metals, notably iron and manganese, with iron often employed as a marker for undesirable concentrations. A disadvantage of the most commonly employed filter aid, kieselguhr or diatomaceous earth, is that impurities in the silica structure solubilise in beer. Although the increase in dissolved iron is small, this can easily be the source of half of the concentration in the beer. The elimination of this effect by the absence of kieselguhr may be of significant benefit to product quality. Reduction in free radical formation through the employment of membranes instead of kieselguhr has been measured (Broens et al., 2007). Processing costs are generally less sensitive to the filterability of the beer for membrane filtration than filter aid filtration. The relative compactness of the plant means that cleaning is more rapid and there is less loss of product. Water usage is reduced, although the need to regenerate fouled membranes means that there is increased usage of detergents. There is scope to employ the membranes as sterile filters, negating the need for pasteurisation or secondary sterile filtration. However, current systems are mostly employed only as primary filters. This is because many brewers employ stabilisation processes downstream of filtration. Also, sterile filtration would mean that each module would require integrity testing (confirmation that the membrane pores remain small enough for sterilisation) before each filter run. There is also the possibility that membranes could fail midrun, although there is instrumentation that employs advanced light-scatter techniques that are claimed to detect membrane failure (Wilhelm, 2009). There is likely to be enormous potential for development of improved membrane process efficiency. For example, membrane manufacturing techniques derived from the electronics industry have been employed to manufacture so-called microsieves (Lommi et al., 2003). These comprise silicon nitride wafers. They feature a very high porosity, very precise and consistent pore diameter and membrane thickness of the order of just 1 mm.
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446 Separation, extraction and concentration processes This enables a requirement for very low trans-membrane pressures, and also the application of a high frequency backflush (‘backpulse’) to maintain permeate flow rate. Consequently, the membranes achieve a permeate flow rate per unit area with beer of the order of a hundred times greater than with PES membranes. However, the current absence of mass production that would allow economical full-scale plants, and also problems optimising the backpulse in larger plants, have as yet prevented successful commercialisation of the technology.
15.11 Recovery of cleaning detergents in brewing Operation of modern breweries is characterised by the application of cleaningin-place (CIP) technology. This is characterised by a cycle of recirculating water rinses, detergents and disinfectants in the appropriate sequence. Advantages include effective automation, reliable programmable cleaning cycles and the ability to maintain a closed process. Water economy is achieved by reusing final rinses as first rinses and automation of the disposal of spent cleaning agents and automated top-up of cleaning agent strength. Nevertheless, a single typical cleaning cycle employs a further 10% of the water that finishes in the product (Freeman, 2008). However, application of nanofiltration membrane technology enables removal of fouling from the detergent. The detergent, either acidic or alkaline, is recovered at the rate of typically 90%. It is mostly the relatively small detergent molecules that pass through the membrane, but there is some transmission of soil so that some system purging is required. The economic viability of the installation is improved with economy of scale. As with other membrane processes, viability is likely to improve in the future as membrane performance is improved and membrane prices fall in real terms. Also the prices of water, water treatment and detergents are certain to increase in the future. Currently, dependent on process conditions and geographical location, payback time for capital expenditure has been calculated as less than two years (Catala and Freeman, 2008). A more recent calculation suggests that payback time will actually be typically less than one year (Freeman, 2008). Future improvement to the technology could be automation based on in-line measurement of contamination in the detergent, perhaps based on absorption of near infra-red radiation.
15.12 Dissolved gas control by membrane technology As a carbonated beverage, the concentration of carbon dioxide in beer is one of the most important factors contributing to its flavour. In a modern brewery, most typically, carbon dioxide has to be added to attain the relatively high
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Separation technologies in brewing 447 specification of most small-pack or keg beers. This is often achieved by inline post-filtration, with a feedback control loop and controlled dosage of carbon dioxide into a turbulent zone that ensures dissolution. However, on occasions there is a need to reduce carbon dioxide concentration to bring it within the specification. This is problematical, because it usually requires a nitrogen purge that risks loss of foam potential in the product and can form unsightly particles as the foam collapses. Furthermore, the need to reduce carbon dioxide concentration may also occur because the beer is a mixed gas or ‘nitro-keg’ product. These products exhibit a ‘smooth’ palate, enhanced foam potential with an attractive ‘theatre in dispense’ (Cooper, 2009). Thus there is scope for a process that can adjust gas concentrations in beer without dispersing bubbles. A suitable technology relies on hydrophobic membranes (Gill and Menneer, 1997). These comprise hollow membrane fibres made of a hydrophobic polymer, commonly polypropylene. The membrane pores do not allow passage of the beer up to a certain pressure. To achieve the pressure differential required in brewery processing, the pores have to be as small as 50 nm in diameter. The beer is circulated on one side of the fibre (most probably the tube side) and a particular mixture and pressure of gases is applied to the other side. A vacuum will allow a reduction of carbonation. Application of nitrogen allows nitrogenation of the beer whilst simultaneously reducing carbonation, because each gas diffuses according to its own partial pressure. A further advantage is that because there is no oxygen in the applied gases, there is potential to reduce the residual oxygen content of the beer, assisting product stability. The technology has been successfully applied in UK breweries making ‘nitro-keg’ beers (Cooper, 2009).
15.13 Future trends As with all industries, future business strategies are required to pay greater attention to environmental issues such as water and energy economy and waste reduction. There is the prospect of standardisation of calculation of water and energy footprints that will become increasingly important key performance indicators. It appears that the brewing industry is partitioning into two distinctive types of company: major international brewing companies that market ‘key brands’ globally and local concerns that produce perhaps more ‘niche’ products. In the former instance especially, there is scope for employing key performance indicators as marketing tools to gain advantage over brand rivals. It is possible that there is a significant move towards beer products of relatively low alcoholic strength (Evans, 2008). This would enable beer to be produced by still brewing to high alcoholic strength (‘high gravity brewing’) but increasing rates of dilution near the end of the process, thus economising on utility usage. © Woodhead Publishing Limited, 2010
448 Separation, extraction and concentration processes Rising utility and effluent treatment costs may result in more widespread introduction of sophisticated processing of effluent streams. For example, anaerobic digestion may be employed to generate biogas fuel from effluent streams. The residual effluent could be further sequentially processed culminating in reverse osmosis. The resultant water is effectively completely clean, nearly pure, and could be reused as a brewing raw material.
15.14 References Anon. (2004), Official inauguration of kieseguhr-free filter line at Pott’s Brewery in Oelde/Westphalia (Germany), Brauwelt International, 22(5), 372–373. Anon. (2009), Greenfield brewery with kieselguhr free filtration technology, Filtration, 9(3), 169. Bamforth C W (2000), Making sense of flavour change in beer, Technical Quarterly of the Master Brewers’ Association of the Americas, 37(2), 165–171. Borremans E and Modrok A (2003), Membrane filtration by Alfa Laval and Sartorius, Brewer and Distiller International, 34(4), 10–11. Boulton C (2009), Stirring stuff. Getting the best out of cylindroconical fermenters, Brewer and Distiller International, 5(6), 18–21. Broens L, Meijer D, Mepschen A, Schuurman R, Methner F, Kunz T, Eisenblatter F, Metz L and Brunacker J (2007), Practical membrane filtration for beer clarification, Proceedings of the European Brewery Convention, Venice. Brocheton S, Hermia J, Rahier G and van den Eynde E (1995), The basic principles of a new beer filtration process, Proceedings of the European Brewery Convention Brussels, Nurnberg, Fachverlag Hans Carl, 427–436. Browne J (2008), Innovation is about survival. The challenges of a modern small pack line, Brewer and Distiller, 4(1), 11–17. Brunner M (1987), Filter sheets, International Bottler and Packer, 61(6), 27–32. Catala M and Freeman G (2008), The relationship between water consumption and energy usage in the malting and brewing industries: opportunities and priorities, Proceedings of the World Brewing Congress 2008. Cooper D (2009), The nitro-keg revolution, Beers of the World, 24, 54–56. Davies K (2004), Filter aids, The Brewer International, 4(2), 14–18. Duan D, Rogers P, Dawson J, Aspridis C, Day G, Delaere S and Oliver T (2008), The use of pectin-based finings in commercial-scale beer making, Proceedings of the 30th convention of the Asia Pacific Section of the Institute of Brewing and Distilling, Auckland. Evans E and Kaur M (2009), Keeping sleepy yeast awake until bedtime, Brewer and Distiller International, 5(5), 38–40. Evans J (2008), Low alcohol beers reach new heights, Brewers’ Guardian, 137(10), 29–32. Eyben D, Hermia J, Meurens J, Rahier G and Tigel R (1989), Industrial results of a new wort filter, Proceedings of the 22nd Congress of the European Brewery Convention, Zurich, 275–281. Ferstl FF and Zuber J (2009), Pre-coat filtration with a new, regenerable filter aid, Proceedings of the 13th Scientific and Technical Convention of the Institute of Brewing and Distilling Africa Section. Freeman G (2008), Cleaning-in-place, Campden BRI’s ‘Cleaning and disinfection conference – managing new challenges’. Freeman G J, Powell-Evans M H B, Baron J M, Dawson M K, Patel A, Skipper A J, Evans C T, Boulton C A, Grimmett C M and Le Gourrierec X (2003), Improving the © Woodhead Publishing Limited, 2010
Separation technologies in brewing 449 effectiveness of isinglass finings for beer clarification by optimisation of the mixing process part 3: full-size prototype evaluation, Journal of the Institute of Brewing, 109(4), 326–331. Gill C B and Menneer I D (1997), Advances in gas control technology in the brewery, The Brewer, 83(987), 77–84. Harmegnies F, Bonacchelli B and Tigel R (1997), Beer filtration with regenerable filter aid, Proceedings of the European Brewery Convention Maastricht, Nurnberg, Fachverlag Hans Carl, 517–524. Harmegnies F, Bonacchelli B and Tigel R (2003), Continuous brewing, Proceedings of the 29th Congress of the European Brewery Convention, Dublin. Lake J C and Speers R A (2008), A discussion of malt-induced premature yeast flocculation, Technical Quarterly of the Master Brewers Association of the Americas, 45(3), 253–262. Lommi H, Raspe O J, van Rijn C J M and Vos J (2003), New filter and method for beer clarification, Proceedings of the 29th European Brewery Convention Congress. Malfliet S, Van Opstaele F, De Clippeleer J, Syryn E, Goris K, De Cooman L and Aerts G (2008), Flavour instability of pale lager beers: determination of analytical markers in relation to sensory ageing, Journal of the Institute of Brewing, 114(2), 180–192. Martin I (2002), Pasteurisation possibilities for plastics, Brewers’ Guardian, 131(1), 20–22. Meckler O (2003), New generation centrifuges in breweries, Proceedings of the 9th Brewing Convention of the Institute and Guild of Brewing Africa Section, Victoria Falls, Zambia, 178. Noordman T R, Berghuis O A E, Mol M N M, Peet C J, Muller J L M, Broens L and van Hoof S (1999), Membrane filtration for bright beer, an alternative to kieselguhr filtration, Proceedings of the 27th Congress of the European Brewery Convention, 815–822. O’Rourke T (1999), Mash separation: a review, Brewers’ Guardian, 128(2), 15–16. O’Rourke T (2000), Microbiological quality, Brewers’ Guardian, 129(10), 39–42. Pickerell A and Heeb W (2008), Cross-flow membrane filtration at the Coors Shenandoah brewery, World Brewing Congress 2008 Proceedings. Poku M (2004), An investigation into the recycling of filter aids for the brewing industry, PhD Thesis, University of Essex Department of Biological Sciences. Riddell P (2002), Sterile filtration of beer. The PES story, Brewer International, 2(11), 31–35. Russ W (1993), Disposal of kieselguhr – kieselguhr recycling, Brauwelt International, 11(1), 51–55. Siebert K J (1997), Beer clarity stability, Proceedings of the 6th Central and Southern Africa Section Convention of the Institute of Brewing, 67–78. Soares E V (2009), Flocculation in Saccharomyces cerevisiae, in Preedy V R, Beer in health and disease prevention, London, Elsevier, 103–112. Tubbs J (1998), Cartridge filtration – part 2, The New Brewer, 15(3), 56–58. Virden J (1995), How to avoid ‘mincing’ your trub, BDI, 26(11), 20–21. Waiblinger R (2002), Beer filterability, The Brewer International, 2(1), 15–18. Wilhelm M (2009), Continuous integrity monitoring of membrane filter units, Brauwelt International, 27(3), 143–145. Zimmerman T and Meffert M (2008), Crosspure® – The future of kieselguhr-free beer filtration, Proceedings of the 30th Convention of the Asia Pacific Section of the Institute of Brewing and Distilling, 1 p.
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450 Separation, extraction and concentration processes
16 Methods for purification of dairy nutraceuticals C. J. Fee, J. M. Billakanti and S. M. Saufi, University of Canterbury, New Zealand
Abstract: The properties of the classes of proteins found in bovine milk and increasingly used as nutraceuticals are reviewed. Whey proteins, an important class of dairy nutraceutical products, can be classified as acidic and basic proteins and immunoglobulins. The methods used for their purification, including ion exchange, chromatography and membranes are described. Key words: nutraceuticals, milk proteins, lactoferrin, ion exchange, chromatography, membranes.
16.1 Introduction The global market for nutraceuticals was worth US$117bn in 2007, US$124bn in 2008 and is expected to reach US$177bn by 2013 at a compound growth rate of 7.4% per year, driven in part by increased demand from developing countries. Many biologically active components and derivatives of milk and their nutraceutical applications have been reviewed by Severin and Xia (2005). These include minor milk proteins such as lactoferrin, lactoperoxidase, immunoglobulins and lysozyme, endogenous peptides and those derived from protein hydrolysis, oligosaccharides, hormones, growth factors and gangliosides. Of the proteins, it is mainly the whey proteins that have nutraceutical applications. However, casein micelles have been proposed as a novel delivery vehicle for nutraceutical compounds, taking advantage of the casein protein self-assembly (Semo et al. 2006). In addition, milk proteins, including caseins, are a source of biologically active peptides that are inactive within the sequence of a native protein but can be released by enzymatic hydrolysis.
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Methods for purification of dairy nutraceuticals 451 Such bioactive peptides have been found to exhibit various physiological activities such as antihypertensive, opioid, immunomodulatory antimicrobial, antioxidative, antithrombotic, and cytomodulatory activities and may find use in the treatment of diarrhea, hypertension, thrombosis, dental caries, oxidative stress, mineral malabsorption, and immunodeficiency (Haque et al. 2009). Lactose, the main carbohydrate in milk, can be extracted and crystallized, and its derivatives used as nutraceuticals (Ganzle et al. 2006). However, because enzymatic hydrolysis is generally required to produce bioactive peptides and chemical modification is required to produce bioactive lactose derivatives, rather than extraction/purification of either class of nutraceuticals in their native form, we will not consider these products further here. The milk fat globule membrane has recently been shown to contain a rich variety of components that have an impressive array of functionalities, including cholesterolemia-lowering factor, inhibitors of cancer cell growth, vitamin binders, inhibitor of Helicobacter pylori, inhibitor of b-glucuronidase of the intestinal Escherichia coli, xanthine oxidase as a bactericidal agent, butyrophilin as a possible suppressor of multiple sclerosis, and phospholipids as agents against colon cancer, gastrointestinal pathogens, Alzheimer’s disease, depression, and stress (Spitsberg 2005). However, these components are present at extremely low levels, so their economic fractionation would be challenging. The basic whey proteins, lactoferrin, lactoperoxidase and lysozyme, have known antimicrobial properties, whereas the main acidic whey proteins have functions that include lactose synthesis, Ca2+ transport, immunomodulation and anticarcinogenicity (a-lactalbumin) and retinol transport, fatty acid binding and antioxidant activity (b-lactoglobulin). Lactoferrin, in particular, has multiple activities, including antioxidant activity, iron binding, anticarcinogenicity and immunomodulation (Severin and Xia 2005). Immunoglobulins, unsurprisingly, impart various immunoprotective functions in a range of functional foods (Gapper et al. 2007). Table 16.1 shows the main components of milk, their concentrations and their bioactivities relevant to nutraceuticals (Severin and Xia 2005). In this chapter, we focus on the purification of whey proteins as nutraceuticals. The main workhorses of whey protein purification are ionexchange chromatography and membrane separations. However, examples of affinity chromatography, adsorptive membranes and other separation formats aimed at large-scale processing are also presented. We have organized the sections by product (acidic proteins, basic proteins and immunoglobulins) and describe examples for the separation/purification of each.
16.2 Components of acidic whey protein 16.2.1 b-Lactoglobulin: properties and applications b-Lactoglobulin (b-Lac) is the major protein component in bovine whey, constituting approximately 58% of the whey protein or 10% of total milk © Woodhead Publishing Limited, 2010
452 Separation, extraction and concentration processes Table 16.1 Milk proteins and their bioactivities (adapted from Severin and Xia, 2005) Concentration (g L–1) Protein Total caseins
Human
Cow
2.7
26.0
a-Casein
13.0
b-Casein
9.3
k-Casein
3.3
Total whey protein
6.3
Ion carrier, bioactive peptides precursors
6.3 3.2
b-Lactoglobulin
Bioactivity
a-Lactalbumin
1.9
1.2
Immunoglobulins (A, M, G) Serum albumin
1.3
0.7
0.4
0.4
Lactoferrin
1.5
0.1
Lactoperoxidase Lysozyme
0.1
Miscellaneous
1.1
Retinol carrier, fatty acid binding, possible antioxidant Lactose synthesis, Ca2+ carrier, immunomodulation, anticarcinogenic Immune protection
Antimicrobial, antioxidative, immunomodulation, iron adsorption, anticarcinogenic 0.03 Antimicrobial 0.0004 Antimicrobial, synergies with immunoglobulins and lactoferrin 0.8
Proteose-peptone
1.2
Glycomacropeptide
1.2
Antiviral, bifidogenic
protein (Lozano et al. 2008). The concentration of b-Lac in whey is in the range 2–4 g L–1 (Andersson and Mattiasson 2006). The primary structure of b-Lac consists of 162 amino acids and it has a molecular weight (MW) of approximately 18.4 kDa. Six genetic variants are known, the most common being the A and B variants. The A and B variants differ in only two amino acid residues, at positions 64 and 118, which are Asp and Val for b-Lac A and Gly and Ala for b-Lac B (Elofsson et al. 1997). Because b-Lac A has an additional negative charge, it has slightly lower pI value (pI 5.1) than b-Lac B (pI 5.2), although the molecular weights of the variants are essentially the same (Yamamoto and Ishihara 1999). The secondary structure of b-Lac comprises nine strands of b structure, a short a helix segment and three helical turns. Its quaternary structure depends on the medium pH: it occurs mainly as a stable dimer, with a molecular weight of 36.7 kDa, at pH values between 5.2 and 7; as an octamer, with © Woodhead Publishing Limited, 2010
Methods for purification of dairy nutraceuticals 453 a molecular weight of 140 kDa, at pH values between 3.5 and 5.2; and as a monomer, with two-cysteine residues per monomer, at pH below 3.0 and above 8.0 (de Wit 1989). b-Lac is a good source of essential amino acids and has a potential use in power drinks owing to its good solubility (Horton 1995). Good gelling formation and superior foam stability compared with other whey proteins make b-Lac suitable for confectionery production (Cowan and Ritchie 2007; Zydney 1998). Madureira et al. (2007) summarized a number of biological functions of b-Lac: it plays a role in regulation of mammary gland phosphorus metabolism, as a transporter for vitamin D, cholesterol and retinol, the transfer of passive immunity to the newborn and enhancement of pregastric esterase activity (Madureira et al. 2007). The b-lac content of bovine milk is much higher than that of human milk (El-Agamy 2007; Fox and McSweeney 1998) and this has been identified as a potential source of allergic reactions to infant formulae seen in some children (El-Agamy 2007; Monaci et al. 2006; Suutari et al. 2006). Therefore, the removal of b-Lac from whey may broaden the applications of whey product derivatives in the food industry (Casal et al. 2006). b-Lac-free whey may serve as the primary protein constituent of hypoallergenic infant formulae that have protein compositions that are more similar to that of human milk (Casal et al. 2006). 16.2.2 a-Lactalbumin: properties and applications a-Lactalbumin (a-Lac) is the second largest protein component of bovine whey, comprising approximately 3.4% of total milk protein or 20% of whey proteins. The concentration of a-Lac in whey protein is 1.2–1.5 g L–1 (Andersson and Mattiasson 2006). On the other hand, a-Lac is the predominant whey protein in human milk, with a concentration of 2.44 ± 0.64 g L–1 (after day 30 of lactation) (Jackson et al. 2004). a-Lac is a strong Ca2+-binding protein, consisting of 123 amino acids in a single peptide chain with four disulfide bonds. Its molecular weight is about 14.2 kDa and it has a pI value of 4.2. Human and bovine a-Lac have approximately 76% fully conserved residues (93 out of 123 amino acids) (Chatterton et al. 2006). a-Lac is the preferred protein source for infant formulae, owing to its high tryptophan content, high digestibility and lower potential for causing allergies compared with b-Lac (Gurgel et al. 2000; Zydney 1998). Additionally, because of its high tryptophan content it is applicable as a nutraceutical and because of its high cytotoxicity it possesses therapeutic uses (Konrad and Kleinschmidt 2008). a-Lac also has strong affinity for glycosylated receptors on the surface of oocytes and spermatozoids and may thus have potential as a contraceptive agent (Zydney 1998). Its reported biological functions include cancer prevention, lactose synthesis and treatment of chronic stress-induced diseases (Madureira et al. 2007).
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454 Separation, extraction and concentration processes 16.2.3 Bovine serum albumin: properties and applications Bovine serum albumin (BSA) exists in whey at concentrations between 0.3 and 0.6 g L–1, its molecular weight is about 69 kDa and it has a pI value of around 4.7–4.9 (Andersson and Mattiasson 2006). It consists of 582 amino acid residues, with 17 intramolecular disulfide bonds and a single free thiol at residue 34 (Burr 2001). Biological functions of BSA include fatty-acid binding, an antimutagenic function and cancer prevention (Madureira et al. 2007). BSA also has good gelling properties (Matsudomi et al. 1991) and is widely used in food and therapeutic applications (Zydney 1998).
16.3 Purification technologies for acidic whey proteins 16.3.1 b-Lactoglobulin purification Extraction of b-Lac from whey has been achieved by various technologies, including chromatography, membrane filtration, selective chemical and thermal precipitation, membrane chromatography, foam fractionation and colloidal gas aphrons. A number of recent examples of acidic protein purification are described briefly here, with many further examples from the past 10–15 years tabulated in Tables 16.2 to 16.5. Selective precipitation As described below, b-Lac has been selectively isolated from whey by forming a complex with chitosan (Casal et al. 2006; Montilla et al. 2007), addition of ammonium sulfate (Lozano et al. 2008), precipitation of a-Lac with sodium citrate (Alomirah and Alli 2004) and by peptic hydrolysis followed by membrane filtration (Konrad et al. 2000). b-Lac interacted reversibly with chitosan by electrostatic interaction, forming a precipitate at pH 6.2. The b-Lac was recovered by dissolving the precipitate in 100 mM sodium acetate at pH 9 to give a recovery of 90%, with a purity of 95% (Montilla et al. 2007). The isolated b-Lac maintained its native structure and the use of non-toxic chitosan may be of interest in industrial applications. b-Lac was separated from other whey proteins by precipitation with 50% ammonium sulfate (Lozano et al. 2008). The precipitate was dissolved and then again separated using 70% ammonium sulfate, leaving a supernatant solution enriched in b-Lac. After dialysis, lyophilization and reconstitution in water, final purification is carried out by weak cation-exchange chromatography. The total yield and purity of b-Lac from 3.5 L whey were 14.32 and 95%, respectively. Meanwhile, Alomirah and Alli (2004) recovered b-Lac from supernatant solution after a-Lac precipitation with sodium citrate. After several further steps (i.e. washing, centrifuge and dialysis), b-Lac was recovered with a purity ranging from 83 to 90%. The yield of the b-Lac isolate from this process was reported to be in the range 47–69%. © Woodhead Publishing Limited, 2010
Methods for purification of dairy nutraceuticals 455 Table 16.2 Additional examples of membrane separations of acidic whey proteins Year
Author
Target Protein source protein
2009 Metsamuuronen a-Lac and Nystrom b-Lac
Configuration/material
Whey powder and fresh whey
Flat sheet (2 ¥ 10–3 m2): 30 kDa and 100 kDa regenerated cellulose, 50 kDa polyaramide, 20, 30 and 50 kDa PES, 100 kDa PSF Flat sheet: 30 kDa regenerated Single 2009 Bhushan and b-Lac, cellulose (Millipore), 25 mm Etzel b-Lac, single GMP glycomacropeptide diameter (GMP), binary b-Lac and GMP, whey 2008 Konrad and Flat sheet: 100 kDa PES (0.093 a-Lac Concentrated Kleinschmidt rennet whey m2). Spiral wound: 150 kDa PSF (5.5 m2) 2007 Cowan and Flat sheet: 100 kDa PES (47 mm a-Lac, Single a-Lac, Ritchie diameter) b-Lac single b-Lac 2007 Almecija et al. a-Lac, Whey Tubular ceramic membrane (Clover Inside Ceram, TAMI b-Lac, Industries, France), 300 kDa, area BSA, 0.045 m2 LF, IgG 2006 Bhattacharjee b-Lac, WPC (obtained Flat sheet: 10 and 30 kDa PES et al. a-Lac after pretreating (56 mm diameter) the whey) 2004 Cheang and b-Lac, WPI (spiked with Flat sheet: 30 and 100 kDa Zydney regenerated cellulose, 50 cm2 a-Lac, some BSA) BSA 2003 Muller et al. 50 kDa Ceram Membrane (TAMI a-Lac WPC (2003a) Industries, France), average pore diameter 12 nm, area 0.045 m2 1998 Lucas et al. Tubular: inorganic membrane a-Lac, WPC (Carbosep, Tech-Sep), 10, 15, 50 b-Lac and 150 kDa, ID 0.6 cm, OD 1 cm, length 60 cm, area 0.0113 m2 WPI, whey protein isolate; WPC, whey protein concentrate; UF, ultrafiltration; MF, microfiltration; PES, polyethersulfone; PSF, polysulfone.
Konrad et al. (2000) compared several techniques for isolation of b-Lac from whey on a large scale. They developed a fractionation method consisting of peptic treatment of 10 000 L of whey, followed by membrane filtration. Three other methods compared were trichloroacetic acid precipitation, a salting out procedure and selective thermal precipitation. The yields of native b-Lac achieved by these four methods were 67.3, 44.9, 46.7 and 49.6%
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© Woodhead Publishing Limited, 2010
Year
Author
Target protein
Protein source
Materials/configuration
Mode of interaction/ligand
2008
Brochier et al.
b-Lac
Microfiltered whey
Mixed mode–hexyl amine
2008
Etzel et al.
WPI
Whey
2008
Etzel et al.
WPI
Whey
2006
Liang et al.
Whey
2004
Schlatterer et al.
b-Lac, a-Lac, BSA, IgG b-Lac
HyperCel™ column (Pall BioSepra), crosslinked cellulose, 90 mm, column volume 2.5, 5 and 10 mL Mono S column (Amersham Pharmacia Biotech Inc.), 2.38 L, 10 cm diameter SP Sepharose Big Beads (Amersham Biosciences), 5.34 L column volume, 20 cm diameter, 17 cm height Sephadex G-200, 2.6 cm ¥ 70 cm
2004
Turhan and Etzel a-Lac, WPI
2004
Rojas et al.
a-Lac, b-Lac
2004
Doultani et al.
2003
Neyestani et al.
a-Lac, WPI, LP, LF b-Lac, a-Lac, BSA
2002 2002
Vyas et al. Lan et al.
b-Lac Total protein
Macro-Prep ceramic hydroxyapatite (BioRad), 80 mm, 12 mm × 88 mm Lactic acid SP Sepharose Big Beads (Amersham whey Biosciences), 80 mL column volume Protein fraction Sephadex G-25® HR-10/10 (Amersham Bioscience) from ATPS Whey SP Sepharose Big Beads (Amersham Bioscience), 80 mL column volume Whey (1) Sephadex G-50 (Amersham Biosciences), 65 cm length, 1.6 cm diameter, 130 mL column volume (2) DEAE column (Amersham), column volume 5 mL Whey Calcium biosilicate particles Whey powder Diaion HPA25 (Sigma)
Cation exchange–methyl sulfonate Cation exchange–sulfopropyl (SP) Gel filtration
Whey
Cation exchange–SP Size exclusion Cation exchange, SP Size exclusion and anion exchange
Affinity, ligand – all-trans-retinal Ion exchange
456 Separation, extraction and concentration processes
Table 16.3 Additional examples of chromatographic separations of acidic whey proteins
2001
Gurgel et al.
a-Lac
WPI
2000
Tellez and Cole
b-Lac, a-Lac, BSA, IgG
Whey
Polyhydroxylated methacrylate (TosoHaas Peptide AF Chelate 650) Biogel A 0.5m and 5m (Bio-Rad, Hercules, Electrochromatography CA, USA), column 1: 1.5 cm diameter ¥ 30 cm height, column 2: 2.5 cm diameter ¥ 60 cm height
ATPS, aqueous two phase system; DEAE, diethyl aminoethyl cellulose.
Methods for purification of dairy nutraceuticals 457
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458 Separation, extraction and concentration processes Table 16.4 Additional examples of adsorptive membrane separation of acidic whey proteins Year
Author
Target protein
Protein source Membrane configuration/ Ligands materials
Whey, single b-Lac Whey, single b-Lac, a-Lac and BSA, binary b-Lac and BSA 2006 Bhattacharjee b-Lac, a-Lac Permeate from et al. two stage UF 2009 Saufi and Fee 2008 Goodall et al.
1996 Splitt et al.
b-Lac, a-Lac, BSA b-Lac, a-Lac, BSA
b-Lac, a-Lac, BSA
Whey
1996 Freitag et al. b-Lac, Whey a-Lac, BSA, IgG
Flat sheet ethylene vinyl Q alcohol Sartobind MA (Sartorius) Q, DIEA
Vivapure Q Mini-H (Vivasciences); 240 ml volume Sartorius cellulose based membrane adsorber: MA Q15 (area 15 cm2), MA Q100 (area 100 cm2), MA D15 (area 15 cm2). Synthetic copolymerbased membrane adsorber (area 1300 cm2) Sartorius cellulose-based membrane adsorber: MA Q15 (area 15 cm2), MA MA S15 (area 15 cm2)
Q Q, DIEA
Q, SP
Q, quaternary ammonium; DIEA, diethylamine; SP, sulfopropyl.
for the peptic treatment method, acid precipitation, salting out and thermal precipitation, respectively. The purity of b-Lac achieved by all methods was more than 90%. Chromatographic techniques Neyestani et al. (2003) used a series of chromatography steps to obtain pure b-Lac from whey after precipitation with 50% ammonium sulfate. Both the precipitate and supernatant solution obtained were dialyzed and lyophilized for further separation by chromatographic methods. A lyophilized precipitate fraction was reconstituted in distilled water and run onto a gel filtration column (Sephadex G-50, 131 mL column volume, 65 cm length) to obtain a first peak containing a mixture of BSA and casein and a second peak of pure b-Lac. The yield of b-Lac was 166 mg based on 50 mL of starting milk. Meanwhile, the lyophilized supernatant was dissolved in water and loaded onto a diethylaminoethyl cellulose (DEAE) anion-exchange column. Stepwise elution was applied to the column, resulting in two separate peaks consisting of a mixture of BSA and a-Lac in the first peak and pure b-Lac
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Methods for purification of dairy nutraceuticals 459 Table 16.5 Additional examples of other separation techniques for acidic whey proteins Year
Author
Technique
Target Protein Protein source
2009
Shea et al.
Foam fractionation
2008
Lozano et al.
a-Lac, b-Lac Spray dried whey powder, WPI Whey b-Lac
2005 2005
Fuda et al. Tolkach et al.
2004
2001
Alomirah and Alli Baumeister et al. Muller et al. Selective precipitation (2003b) Rodrigues et al. Aqueous two-phase system
2000
Konrad et al.
Ammonium sulfate precipitation 2008 Monteiro et al. Aqueous two-phase system a-Lac, b-Lac WPI 2007 Lucena et al. Acid precipitation Sweet whey, a-Lac WPC, WPI 2006– Casal et al.; Selective precipitation with b-Lac Rennet whey 2007 Montilla et al. chitosan
2003 2003
Colloidal gas aphrons Selective thermal denaturation of b-lac Chelating agent precipitation Expanded bed adsorption
Peptic hydrolysis and membrane filtration
b-Lac a-Lac
Whey WPC
a-Lac, b-Lac Whey, WPC, WPI Orotic acid
Whey
a-Lac
WPC
a-Lac, b-Lac Pure a-Lac, b-Lac, WPC Whey b-Lac
WPC, whey protein concentrate; WPI, whey protein isolate.
in the second. The yield of b-Lac in the second peak was estimated to be about 35 mg mL–1. The mixture of BSA and a-Lac was further applied to a gel filtration column to separate pure protein fractions. The yields of BSA and a-Lac, were about 2.3 and 1.1 mg mL–1, respectively, from the resolved peak. Brochier et al. (2008) demonstrated the feasibility of using a mixed-mode chromatography column for isolation of b-Lac from whey, without the necessity for pH or conductivity adjustment using a hexylamine mixed-mode column (HyperCel™, Pall BioSepra, Cergy, France). A smooth scalability from 2.5 to 10 mL column volume (CV) was achieved to extract all b-Lac from five CV of whey loaded into the column. Bound b-Lac was eluted at pH 4, with a purity estimated to be around 75%. Meanwhile, Schlatterer et al. (2004) used a ceramic hydroxyapatite column (Macro-Prep, BioRad, Munich, Germany) to isolate b-Lac from acid whey originating from the milk of healthy and mastitic cows. A single peak of b-Lac could be eluted at a sodium fluoride concentration of 0.6 M. Using
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460 Separation, extraction and concentration processes whey from a healthy cow, the yield of b-Lac was around 50–55%, with a purity of more than 96%. For mastitic whey, the yield of b-Lac was between 18 and 20%, with low purity, contaminated by immunoglobulins, BSA and lactoferrin. Membrane filtration In membrane filtration, b-Lac is normally recovered, together with BSA and immunoglobulin, as a retentate stream during ultrafiltration (UF). UF typically targets a-Lac as a permeant to produce very high purity. Bhattacharjee et al. (2006) used a two-stage UF membrane, followed by anion-exchange membrane chromatography, to produce a b-Lac fraction with 87.6% purity from whey protein concentrate. Bhushan and Etzel (2009) modified a UF membrane to have a positively charged membrane to separate two proteins of similar size, b-Lac and glycomacropeptide (GMP). b-Lac was retained by the charged membrane, whereas GMP selectivity was increased by over 600% compared with the uncharged membrane. Membrane chromatography Initial use of membrane chromatography for whey protein fractionation was carried out by Splitt et al. (1996) and Freitag et al. (1996). Splitt et al. (1996) demonstrated that chromatographic conditions were transferable from a cellulose- to a polymer-based membrane adsorber carrying the same functional groups for whey protein fractionation. In a two-step salt gradient, a-Lac was eluted at 0.1 M NaCl and a mixture of BSA and b-Lac were eluted at 0.5 M NaCl. By passing 5 L of feed (0.065 mg mL–1 a-Lac and b-Lac, 1 mg mL–1 BSA) through 1300 cm2 of total membrane adsorber area, 116 mg a-Lac was eluted in the first peak and 132 mg of a mixture of b-Lac and BSA was eluted in the second peak. Freitag et al. (1996) investigated the concept of mixed-mode interactions to bind all whey proteins in a single pass of whey through a membrane chromatography column. Two modules of MA Q15 and one module of MA S15 were connected in series and whey was passed through at pH 6. However, elution of the anion and cation modules was done separately, because a-Lac and IgG elute at the same salt concentration. MA Q15 produced a single peak containing a-Lac, BSA, b-Lac A and b-Lac B, whereas MA S15 gave a single peak of IgG. Recent studies by Saufi and Fee (2009) and Goodall et al. (2008) used anion exchange membrane chromatography to selectively bind b-Lac from whey. Goodall et al. (2008) observed that when the anionic membrane was saturated with whey, b-Lac could displace other bound protein from the membrane. This can produce a flow through fraction that is depleted in b-Lac, with concentrations of a-Lac and BSA doubled from their original concentrations in whey.
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Methods for purification of dairy nutraceuticals 461 16.3.2 Purification of a-lactalbumin Membrane filtration A novel variant of UF, known as high-performance tangential flow filtration (HPTFF), has been widely used in separation of a-Lac and b-Lac from whey. HPTFF exploits a number of different strategies to achieve high-resolution separations between proteins of similar size, by selecting specific conditions such as (Van Reis et al. 1997; 1999): ∑
proper choice of pH and ionic strength to maximize differences in the hydrodynamic volumes of the product and impurities, ∑ use of electrically charged membranes to enhance the retention of likecharged proteins, ∑ operation in the pressure-dependent regime to maximize selectivity, and ∑ use of a diafiltration mode to wash impurities through the membrane. A dramatic improvement in a-Lac permeation and selectivity can be achieved, as demonstrated by Cowan and Ritchie (2007), who obtained five times better selectivity of a-Lac after modifying a PES membrane with a positively charged group. Similar improvement was also shown by Lucas et al. (1998), who used an inorganic membrane coated with positively charged polyethyleneimine. The transmission of b-Lac was reduced to 1%, whereas a-Lac transmission was 10%, giving a selectivity close to 10. Chromatographic techniques Turhan and Etzel (2004) used an SP Sepharose Big Beads (Amersham Biosciences, Uppsala, Sweden) column to isolate a-Lac from lactic acid whey, achieving a purity of 93%. Gurgel et al. (2001) used a hexamer peptide ligand to bind a-Lac from whey protein isolates, attaining a purity of about 87%, with lactoferrin being the main impurity. Doultani et al. (2004) used a selective elution method to recover different fractions of whey protein bound onto a cation-exchange column at pH 4. Different elution solutions could be applied to produce: (1) a whey protein isolate (WPI), using 10 mM NaOH, (2) a-Lac, using 100 sodium acetate at pH 4.9 and WPI depleted in a-Lac, using 10 mM NaOH, (3) a-Lac, using 100 sodium acetate at pH 4.9, WPI depleted in a-Lac, using 50 mM sodium phosphate at pH 6.5, lactoperoxidase, using 0.35 M NaCl in 50 mM sodium phosphate at pH 6.5 and lactoferrin, using 1.20 M NaCl in 50 mM sodium phosphate at pH 6.5. According to the authors, this method offers the flexibility to switch between different protein fractions, day-to-day, depending on the market and customer demands.
Other techniques Chemical precipitation with sodium hexametaphosphate (Alomirah and Alli 2004) was used to recover a-Lac from whey. The yield of a-Lac was reported to be between 44 and 89%, with a-Lac purities between 86 and
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462 Separation, extraction and concentration processes 90% (Alomirah and Alli 2004). Tolkach et al. (2005) used selective thermal precipitation to isolate native a-Lac from whey protein concentrate. Before precipitation, the environmental parameters of WPC were optimized in terms of total protein, lactose and calcium content, and pH value. The purity of a-Lac achieved by selective thermal precipitation of optimized properties of WPC was 98%, with a recovery about 75%. Muller et al. (2003a, 2003b) used a two-step process to purify a-Lac from whey protein concentrates. In the first step, whey was filtered through a 30 kDa UF membrane with operating conditions that enhanced the ratio of a-Lac/b-Lac in the permeate stream by minimizing the passage of other whey proteins (Muller et al. 2003b). In the second step, two options were investigated, a second UF module or a selective precipitation route (Muller et al. 2003b). The precipitation route was more promising compared with UF, with the purity of a-Lac achieved in the range of 77–99% and yields of 46–83%, depending on the permeate properties from the first UF step. 16.3.3 Purification of bovine serum albumin Bovine serum albumin (BSA) is not commonly isolated as a specific product from whey but rather comes about as a by-product in a-Lac or b-Lac extraction methods. In UF, BSA is normally recovered in the retentate stream with b-Lac and immunoglobulins. In chromatographic separations, BSA typically binds to anion-exchange columns, but, by appropriate selection of elution buffers, BSA contamination in b-Lac or a-Lac fractions can be minimized.
16.4 Basic proteins in the dairy nutraceutical industry The most important basic proteins present in milk are lactoferrin (LF) and lactoperoxidase (LP). Lysozyme is also present in very minor quantities but it is not commercially viable to process this protein from milk. The concentration of LF, LP and lysozyme in milk varies from species to species, breed, stage of lactation, parturition, nutrition, udder health and season of lactation (Thomson et al. 2005). Below are some of the physicochemical characteristics of the basic whey proteins. 16.4.1 Lactoferrin Lactoferrin (LF) is a well-characterized iron-binding glycoprotein that belongs to the transferrin family, also known as lactotransferrin (LTF). The nonglycosylated form of LF has a molecular weight of 80 kDa (690 residues) and a pI of 8.6. LF is present in several secretory substances, including milk, tears and saliva (Masson et al. 1966). It is one of the major whey proteins in human milk, with a concentration of about 1.4–2.0 mg mL–1 but it is only a
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Methods for purification of dairy nutraceuticals 463 minor component in bovine milk, at only one-tenth of these concentrations (0.1–0.2 mg mL–1). Huge variations in lactoferrin concentrations (0.06–1.0 mg mL–1) of milk from individual cows has been monitored and reported in the literature (Indyk and Filonzi 2005). However, elevated levels of LF can be produced in bovine milk with the help of recombinant technologies (Thomassen et al. 2005). 16.4.2 Lactoperoxidase Lactoperoxidase (LP) was first identified and reported as an endogenous enzyme in milk by Arnold (1881). It has a single polypeptide chain with a molecular weight of approximately 78 kDa with 612 amino acid residues and a pI of 9.6 (Seifu et al. 2005). LP is one of the most abundant and heat-stable enzymes in the milk of many mammals and bovine milk contains about 20 times higher levels than that of human milk (Gothefors and Marklund 1975) and is the second most abundant enzyme in bovine milk. Concentrations of LP and its catalysing activity in milk depends on several factors, including feed, stage of lactation and the breed and health status of the animals (Fox and Kelly 2006). A typical concentration of LP was reported as 30–40 mg L–1 (Seifu et al. 2005) in bovine milk. The protein is less important than LF in terms of nutraceutical products but it is crucial in milk preservation and storage and binds to cation exchangers during isolation of LF. 16.4.3 Lysozyme Lysozyme is one of the most abundant enzymes present in human milk. Human lysozyme contains 130 amino acid residues and has a molecular mass of 14.7 kDa and a pI of 11.4. Lysozyme is present in a number of secretions, such as tears, saliva, urine, mucus and milk. Chicken egg white is the richest source of C-type lysozyme, having a concentration ranging between 3.4 and 5.8 g L–1 (Wilcox and Cole 1957; Sauter and Montoure 1972) and this protein is very closely related to a-lactalbumin in both sequence and structure. Although all mammals contain C-type lysozymes, they vary widely in terms of structure and physicochemical properties, such as folding/unfolding, structure, calcium binding, stability to heat and pH and pI. Among all mammalian species, ass (1428 mg L–1), mare (790–1330 mg L–1) and human milk (270–890 mg L–1) are the top three sources of lysozyme but bovine milk contains only a tiny amount (0.05–0.21 mg L–1) (Benkerroum 2008).
16.5 Purification technologies for basic whey proteins in the dairy nutraceutical industry Extraction and isolation of LF in both laboratory- and industrial-scale applications have been achieved using various processing technologies © Woodhead Publishing Limited, 2010
464 Separation, extraction and concentration processes such as ion-exchange, membrane-adsorption, size-exclusion, affinity and hydrophobic interaction-chromatography. The following section describes various separation technologies and their potential applications in the dairy industry. A number of chromatographic and membrane methods used for isolation of LF from various sources is given in Table 16.6. 16.5.1 Lactoferrin and lactoperoxidase purification Ion-exchange chromatography Amongst the basic proteins, LF has been extensively investigated in both laboratory and industrial-scale extraction technologies. However, in most cases LP has been co-eluted as a secondary product because of the similarity in their pI values. The most widely used extraction technology to isolate basic proteins has been cation exchange chromatography. Owing to the complex nature of milk, in general LF extraction is carried out by a sequence of individual processing methods, including casein precipitation, filtration and ionexchange chromatography. Cation-exchange matrices that might be routinely used in basic protein separations include phosphocellulose and carboxymethyl (CM) or sulfopropyl (SP) substituted cellulose, Sephadex and Sepharose. Among these cationic resins, CM-Sepharose (a weak cation exchanger) and SP-Sepharose (a strong cationic exchanger) resins (GE Healthcare, Sweden) are very well characterized for high throughput isolation of basic proteins from milk and whey. During early development, CM-Sepharose resin was the most widely used matrix for isolation of LF from acid and rennet whey at around pH 7.7, but in recent times most LF purification methods investigated have used SP-Sepharose strong cation exchange resins. SP-Sepharose Fast Flow (FF)™ or Big Beads™ and Streamline™-SP resins have been widely used for LF and LP extraction from milk and whey feeds in both lab and industrial-scale applications. Etzel et al. (2000) optimized a chromatographic process using SP-Sepharose Big Beads for isolation of LF from skim milk with >90% purities and 80% recoveries in a single-step packed-bed system. They also optimized Streamline-SP resin as an expanded-bed cation-exchange chromatography medium for isolation of LF and LP from skim milk with >80% recovery at a 200 cm h–1 flow rate. Billakanti and Fee (2009) characterized a cryogel monolith chromatography for extraction of minor proteins (LF) by cation-exchange chromatography from whole milk feeds at 550 cm h–1 in a single-step. Yield and purity of LF extracted using this process were >85 and 90%, respectively. Wu and Xu (2009) reported a purification process for isolation of LF and IgG from bovine colostrum using a serial cation–anion exchange chromatography system with 95 and 97% final purities, respectively. Lu et al. (2007) designed a productionscale technology for extraction of LF from bovine colostrum with final yields of 83% and purities of 94% using an ultrafiltration system coupled with a fast-flow strong cation-exchange chromatography column. Fee and Chand (2006) have investigated and optimized a process to isolate LF and LP from
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Table 16.6 Author and year
Purification examples for the basic whey proteins Configuration/material
Tu et al. 2002 Anti-LF immunoglobulins (IgYLF) immobilized on Sepharose 4B support Ye et al. 2000 Carboxymethyl cation-exchange chromatography
Target protein
Protein source Description of study/results
LF
Methods for purification of dairy nutraceuticals 465
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Colostral whey Final products have purification folds of 103–136 of initial concentration and recoveries of 82–99% Bovine They performed the adsorption process using 50 mM LFa and LFb colostrum phosphate buffer at pH 7.7 LFa was eluted in 0.38 M NaCl and LFb at 0.43 M NaCl buffer solutions. Also investigated variation in the LF concentrations of lactation time and species Li-Chan et al. Immobilized yolk antibody on a LF Milk and Binding capacity was a function of ligand density: at 1998 monoaldehyde-activated support cheese whey 9.2 mg mL–1 of IgYLF, binding capacity was 20% (mol %). This could be increased to 80% by using low ionic strength buffers Camperi et al. Red HE-3B dye affinity chromatography LF Rennet whey This process was achieved with 82% recovery and 98% 2000 purity Jyh-Ping and Microfiltration affinity purification LF and IgG Cheese whey Maximum binding capacity on heparin-Sepharose gel Cheng-Hsin was 124 mg mL–1. LF recovered in this process has 92% activity and 95% purity whereas IgG has 86% activity 1991 and 90% purity Korhonen Review Igs Bovine The latest developments and progress in separation 2004 colostrum, technologies for isolation of Igs in both small and largewhey and milk scale applications Stec et al. Isolation of polyclonal Igs by HPLC Igs Bovine serum Ammonium sulfate precipitation was performed as 2004 a first step of Igs isolation followed by HPLC. This method is applicable for preparative-scale applications DeSilva et al. Conference paper. Novel approaches Whey Dairy whey Recent developments in isolation and purification 2003 to meet the challenges in processing proteins and technologies of various proteins and peptides functional dairy components peptides
Author and year
Continued
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Target protein
Protein source Description of study/results
Xu et al. 2000 Separation of IgG and glycomacropeptides using a polystyrene anion exchanger IRA93 and Amicon YM100 membrane
IgG
Dairy whey
Akita and LiChan 1998
IgG1 and IgG2
Kim and LiChan 1998
Configuration/material
Specific antibody-immobilized immunoaffinity chromatography for isolation of bovine immunoglobulin G subclasses Single-step process for isolation of IgG using avidin-biotinylated IgY chromatography
Konecny et al. Purification of monospecific polyclonal 1994 antibodies from hyper immune bovine whey using immunoaffinity chromatography Labrou and Review Clonis 1994
IgG
Igs
Igs
Successful for concentration of IgG while removing all major whey protein contaminates. A continuous process using IRA 93 resin and 100 kD membranes, designed to produce enriched IgG GMP and WPI ingredients from rennet dairy whey Bovine milk, Gave >98% purity of IgG subclasses. Binding capacities colostrum and of the column were 27 and 38% (molar masses) for whey IgG1 and IgG2, respectively. Antibody columns were stable for more than a year with minimal antibody loss Cheddar Binding capacities achieved in this investigation were cheese whey 50–55% (w/w of IgG and IgY ligands). Bound IgG was eluted at pH 2.8 using glycine–HCl buffer with 99% purity Cheese whey
Various modes of affinity adsorption technologies used for isolation of Igs both in laboratory- and industrialscale applications
466 Separation, extraction and concentration processes
Table 16.6
Methods for purification of dairy nutraceuticals 467 whole milk at milking pH and temperatures using SP Sepharose Big Beads as a packed chromatographic resin. In their investigation, they successfully captured both LP and LF (total dynamic capacity of 49 mg mL–1 resin) from raw milk at 300 cm h–1 flow rate with minor leakage (<5%) in the flow through. Minor proteins captured directly from raw milk were expected to have improved activities and active yields. Andersson and Mattiasson (2006) used simulated moving bed (SMB) technology to isolate LF and LP from whey protein concentrate. SMB technology could be a potential application for large-scale purification processes because it attained 6.5 times the target protein concentration whilst consuming 4.3 times less buffer solution than conventional methods. Noel (2007) has described the largest chromatography column for industrial-scale isolation of LF using expanded-bed-adsorption (EBA) chromatography for processing 200 000 L of crude cheese whey per day. During their optimization process, they passed 85 000 L (three cycles) of crude cheese whey through 950 L of cation-exchange resin at a 900 cm h–1 linear velocity. Bound LF was eluted at mild alkaline conditions with a final product yield of 90–100% and very high purity (with a single band in SDSPAGE). The cation-exchange adsorbent used in this process had a dynamic binding capacity of 27–54 g L–1. Shiozawa et al. (2001) investigated EBA technology using Streamline-SP as a stationary phase for extraction of LF from skim milk with approximately 90% recovery and purity at a 150 cm h –1 linear flow rate. Uchida et al. (1996) obtained a patent for their successful industrial-scale isolation methods to purify LF, secretory components and LP from whey and milk with more than 80% purity of each component using gradient-elution protocols. LP was eluted first, followed by secretory components and then LF. Separated components were biologically active and could be utilized in pharmaceuticals, cosmetics, food supplements and drinks. Kussendrager et al. (1994) claimed a patent for developing an industrial-scale high-throughput chromatography process for isolation of LF and LP from skim milk and whey with more than 80% yields at high superficial velocities (>500 cm h–1). Fee and Chand (2005) reported on-farm capture of high-value minor milk proteins (LF and LP) from the raw milk of individual animals using a fully automated protein fractionation robotic system. This method was rapid, used a single step and was designed to minimize microbial contamination, while offering full product traceability to individual animals. Claycomb (2004) holds a patent for on-farm fractionation of high value milk proteins directly from raw milk using an automated milking system (AMS). Dionysius (1991) investigated a CM-Sephadex resin in a stirred tank system, mixing the resin with whey (cheese, rennet and acid) for 60 min at pH 7.0, and captured more than 80% of LP and 90% of LF, although the resin capacity changed with nature of whey. The bound LF was subsequently eluted by NaCl gradient elution. Heeboll-Nielsen et al. (2004) developed a superparamagnetic cation-
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468 Separation, extraction and concentration processes exchange chromatographic process for fractionation of bovine whey proteins including LF and LP with 28-fold purification over the starting material. The resin could be recovered magnetically in a high-gradient magnetic fishing system, allowing the resin size to be minimized for fast adsorption, while being recovered without the need for fine particle filtration, settling or centrifugation. Ulber et al. (2001) described a two-step membrane processing method for separation of LF from sweet whey, the first membrane removing lipids and other solids and the permeate then passing through a cation exchange membrane adsorber. Plate et al. (2006) used 2-m2 membrane modules for preparative-scale isolation of LF and LP from sweet whey with >90% recovery yields and purities. Chiu and Etzel (1997) described a method for extraction of LF and LP from cheese whey using microporous (3–5 mm) cationic membranes and achieved 50 and 70% recoveries, respectively. Mitchell et al. (1994) used cation-exchange membranes operated in either dead end or cross-flow configuration to isolate LF and LP from cheese whey. Affinity chromatography Affinity chromatography is the second most popular chromatography method for isolation of minor milk proteins from complex protein mixtures on a laboratory scale. Wolman et al. (2007) used Red HE-3B dye coupled with hollow fibers as a membrane affinity chromatography matrix for isolation of LF in a single step, achieving a higher adsorption capacity (111.0 mg mL–1) than the same ligand immobilized on agarose beads (9.3 mg mL–1 resins). They could achieve >94% purity and 91% recovery in the final product using whey and colostrum as feed materials. Chen et al. (2007) evaluated micrometre-sized monodisperse superparamagnetic polyglycidyl methacrylate (PGMA) particles coupled with heparin (PGMA-heparin) as an affinity method for isolation of LF from bovine whey in a single step. This magnetic affinity method resulted in a very high binding capacity of 164 mg g –1 resin and may hold promise as a fast process tool for industrial-scale purification of high-purity LF, although one must always consider the potential for leaching of toxic compounds from affinity ligands. Noppe et al. (2006; 2007) described immobilized affinity peptides as a new platform for rapid development of alternative affinity chromatography ligands for isolation of target proteins from crude feeds. In their method development, they used selected bacteriophages immobilized on a microporous monolith column for extraction of LF from complex fluids such as milk and blood serum. Bound LF was eluted with 1 M NaCl and fractions were recovered with >95% purity. Kawakami et al. (1987) used an immobilized LF-monoclonal antibody resin for a single-step isolation of LF from both human and bovine skim milk and colostrum, with very high purities and recoveries of >97%. Unlike LF, isolation of LP has attracted limited attention because of its low concentration and low commercial value. Shin et al. (2001) used
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Methods for purification of dairy nutraceuticals 469 immunoaffinity chromatography for selective isolation of LP from human whey but such processes have not been extensively studied for large-scale applications. Other chromatographic methods Size-exclusion (gel-filtration) chromatography is not commonly used as an early step in large-volume protein purification because linear velocities are low, but it is often used as a final purification step after other chromatography steps to enhance protein purity. For example, Al-Mashiki and Nakai (1987) applied size-exclusion chromatography for isolation of LF and immunoglobulins from different whey streams. Tomita et al. (2002) reported the possibility of LF and LP isolation from cheese whey or skim milk using a semi-large scale hydrophobic interaction chromatography method. Yoshida (1989) used hydrophobic interaction chromatography followed by DEAE ion exchange chromatography to isolate LF and LP from acid whey. Non-chromatographic methods Noh et al. (2005) selectively purified LF using a cationic surfactant as a micelle foaming agent, manipulating the protein behaviour by changing pH and salt concentration in the aqueous phase and surfactant concentration in the organic phase. LF was partitioned into the aqueous phase whereas all other proteins were solubilized into the organic phase. Fuda (2004) obtained approximately 25-fold enrichment of LF and LP from sweet whey using colloidal aphrons (CGAs), which are surfactant microbubbles generated by intense stirring of the anionic surfactant sodium bis-2ethylhexyl sulfosuccinate (AOT). Noel et al. (2001) investigated a low-cost foam fractionation process as the first step in separating LF from milk and achieved approximately 40% mass recovery at pH 10. Similarly, Saleh and Hossain (2001) developed a semi-batch foaming process to separate LF from a multicomponent mixture of BSA, a-Lac and LF by utilizing protein surface active properties. 16.5.2 Lysozyme purification Owing to its high isoelectric point (pI>11), lysozyme typically co-elutes with other basic proteins (LF and LP), but there have been few studies of the isolation of this protein. However, lysozyme has been successfully purified from milk or acid whey by a combination of affinity chromatography on heparin-Sepharose, followed by gel filtration on Sepharose 4B (Wang and Kloer 1984) or Sephadex-G50 (Boesman-Finkelstein and Finkelstein 1982). Duhaiman (1988) used heparin–Sepharose 4B, Sephadex G-75 and hydroxyapatite for chromatographic extraction of lysozyme from camel milk. Recently, using recombinant technology, human lysozyme was expressed at high concentrations (0.5% dry weight) in transgenic rice seed (Huang et
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470 Separation, extraction and concentration processes al. 2002). Wilken and Nikolov (2006) found cation-exchange purification of lysozyme to be optimal at pH 4.5.
16.6 Immunoglobulins in the dairy nutraceutical industry IgG is an important component of bovine milk and comprises the major protein in bovine colostrum. Consumer acceptance of value-added whey protein products that involve Igs began in Asia and migrated to Europe and the US, targeting sports nutrition, infant formulae, dietary supplements and physiologically functional foods (Gapper et al. 2007). Immunoglobulins are g-globulin proteins that are found in blood and other body fluids of human, bovine and all other lactating species. These are used in the immune system to identify and neutralize bacteria, viruses and other antigens. Human immunoglobulins are broadly classified into five different classes: IgA, IgD, IgE, IgG, and IgM, whereas bovine milk and colostrum mainly contain immunoglobulins IgG, IgM and IgA. All immunoglobulin classes share a basic structural unit that comprises four polypeptide chains with two identical heavy (H) chains and two identical light (L) chains, held together with disulfide bonds. Molecular masses of the heavy- and light-chain peptides are approximately 50–70 kDa and 25 kDa, respectively. 16.6.1 Immunoglobulin G Immunoglobulin G is the most abundant class of antibody in the colostrum and milk of several mammalian species, comprising 80–90% of total antibodies. The primary structure of IgG contains two heavy (each of 450–550 amino acids) and two light chains (each of 211–217 amino acids) with a total molecular mass of approximately 150–160 kDa and a pI of 5.5–8.3. Commonly, IgG presents in two-sub classes such as IgG1 (MW 160 kDa, pI 5.5–6.8) being the richest class, at 15–180 g L–1, and IgG2 (MW 150 kDa, pI 7.5–8.3), at 1–3 g L–1. 16.6.2 Immunoglobulin A Similar to other immunoglobulins, the monomeric structure of IgA also contains two heavy and two light chains but it occurs as a monomer or a dimer, the latter joined by a J-chain and a secretory component. This complex structure is called secretory IgA (sIgA) and has a molecular weight of 380–417 kDa. IgA is present in blood, milk and several mucosal surfaces of lungs and gastrointestinal tracts. This antibody acts as a primary defence system against several pathogens. The IgA concentration in bovine colostrum ranges from 1 to 5 g L–1, whereas bovine milk contains only 0.05 to 1 g L–1.
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Methods for purification of dairy nutraceuticals 471 16.6.3 Immunoglobulin M The monomeric structure of IgM is similar to IgG but IgM consist of five subunits, linked together to form a pentameric structure in a circular mode by disulfide bonds and a J-chain. The molecular mass of IgM is approximately 900–1000 kDa. In bovine milk, the IgM concentration ranges from 0.04 to 1.0 g L–1, whereas bovine colostrum contains 5 to 10 times greater amounts.
16.7 Purification technologies for immunoglobulins in the dairy nutraceutical industry Immunoglobulins from bovine milk have applications as supplements in infant formulae, hyperimmune foods, functional foods, nutraceutical and pharmaceutical products. Although there are many processing methods available for separation of immunoglobulins from various sources, including colostrum and milk, appropriate technologies for large-scale productions are still lacking. Traditionally, isolation of Igs from colostral whey has been achieved by precipitation with either ammonium sulfate or ethanol, followed by chromatography. Although such methods yield rather pure IgG fractions, most are feasible only in small-scale applications. In particular, they are not appropriate for large-scale production of Igs from bovine colostrum or milk (El-Loly 2007). Because milk is a complex fluid and contains high amounts of fat, it is difficult to isolate Igs using conventional methods. However, rapid developments in separation technologies and particularly in the application of membrane separation methods have made industrial-scale isolation of immunoglobulins from various streams, including whey, colostrum and recombinant cell culture supernatants a possibility. 16.7.1 Ion exchange chromatography Use of ion-exchange chromatography to isolate immunoglobulins from dairy feeds is limited but it is used as an intermediate or final purification step in combination with other chromatographic or membrane processing methods. For example, Wu and Xu (2009) applied cation- and anion-exchange chromatography in series for isolation of IgG and LF from bovine colostral whey. Similarly, Pessela et al. (2006) investigated a simple method for isolation of Igs from whey protein concentrate with >80% recovery and reasonable purity using a combination of DEAE–agarose anion-exchange chromatography followed by low-substitution aminated adsorbents. In the first step, they could remove the main contaminant, BSA, and during the second step, Igs were selectively bound on the aminated resin with minor contamination by a-Lac and b-Lac. Noel (2007) developed a second-generation, robust EBA method for
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472 Separation, extraction and concentration processes isolation of Igs from cheese whey, which involved a mixed-mode chemical ligand that has selectivity towards IgG over other whey proteins. The feed whey solution was heated to 50 °C to maintain pasteurization, reduce the viscosity of the solution and maximize flow rate. They could thus load the column at greater than 3000 cm h–1 flow rate without reducing its adsorption capacity. Using a 880-L EBA column, they could produce 13 kg of IgG per day with >90% of final product yields. The eluent from this column was passed through a 160 L anion exchange column for further purification. 16.7.2 Membrane chromatography Since membrane technology was introduced to the dairy industry in the 1970s, many membrane processes have been developed for isolation of Igs from whey and milk using either a single-step process or a combination of several individual steps. For example, UF has been employed either alone or in combination with ion-exchange or gel-filtration chromatography for largescale fractionation of Igs from cheese or colostral whey streams (Abraham 1988; Korhonen 2004; Syväoja et al. 1994). Al-Mashikhi et al. (1988) successfully isolated IgG from acid and cheddar whey using UF, combined with immobilized metal chelate affinity chromatography, obtaining purities of 77 and 53%, respectively. Korhonen (1998) used a multi-step process with UF, microfiltration and reverse osmosis, followed by a cation-exchange resin to achieve 40–75% enrichment of Igs from colostral whey. 16.7.3 Affinity chromatography Affinity adsorption methods have been widely used for isolation of Igs from colostral or cheese whey. Protein-A and Protein-G affinity media are well characterized and used for fractionation of Igs, particularly IgG. Many pilot- and industrial-scale methods using membrane technologies and affinity chromatography have been developed for extraction of Igs from whey and colostrum either alone or in combination with other chromatography techniques (Kochan et al. 1996; Mukkur and Froese 1971; Schmerr et al. 1985). Labrou and Clonis (1994) reviewed various affinity-based purification technologies for Igs. Human and secretory IgA molecules have been isolated by precipitation, ion-exchange and affinity chromatography from human serum (Kondoh et al. 1987; Monteiro et al. 1985; Roque-Barreira et al. 1986) and human milk (Khayam-Bashi et al. 1977). Hutchens et al. (1990) developed a salt-promoted thiophilic adsorption method for simultaneous isolation of all bovine Igs from colostral whey. Mukker and Froese (1971) described isolation of bovine IgM from colostral whey using a combination of molecular sieve and DEAE–cellulose chromatography. Su and Chiang (2003) described a simple non-chromatographic method for isolation of Igs from other whey proteins. They used reverse micelles
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Methods for purification of dairy nutraceuticals 473 to extract more than 90% of IgG into an aqueous phase, with 90% purity, whereas all other whey proteins were partitioned into the reverse micelle phase. 16.7.4 Protein-A mimetic (PAM) affinity peptides Protein A and protein G affinity chromatography can achieve capacities of up to 30–40 mg IgG per mL of resin under optimal conditions. However, these affinity ligands suffer from several inherent drawbacks such as high cost, their biological origin, possible contaminants, low stability towards sanitizing agents and the fact that their selectivity is limited to IgG. Recently, with the help of combinatorial chemistry and modern analytical tools, researchers synthesized protein A mimetic (PAM) peptide ligands (D’Agostino et al. 2008; Fassina 1994; Newcombe et al. 2005; Ulbricht and Yang 2005; Yang et al. 2005) for isolation of human immunoglobulins from mammalian cell cultures. Compared with protein A or G, these synthetic peptide ligands have several benefits (D’Agostino et al. 2008), including high stability towards chemical and biological reagents, low production cost and the absence of biological contaminants. Moreover, these synthetic affinity ligands may also have selectivity towards other classes of immunoglobulins (IgG, IgE, IgM and IgA) (Fassina et al. 1998; Yang et al. 2005), which are not recognized by conventional protein A or G ligands. Several synthetic peptides have been proposed as potential replacements for protein A in affinity chromatography (Ulbricht and Yang 2005). However, the lack of selectivity for antibodies from various classes has limited their widespread use and some of these ligands are not compatible with the retention of antibody activities in proposed buffer solutions. However, more recently, Yang et al. (2005) synthesized a hexamer peptide that has a wide-ranging selectivity towards various classes of immunoglobulins at near-neutral pH. By using a gradient elution with 0.2 M sodium acetate buffer, they were able to isolate various classes of immunoglobulins from human serum (Yang et al. 2005). They have also reported the specificity of this ligand towards immunoglobulins of human, bovine and several other species. Although most of this work has been carried out on human immunoglobulins, there has been recent interest in applying this PAM peptide affinity chromatography for isolation of bovine immunoglobulins from early colostral milk, hyperimmunized cows and transgenic animals.
16.8 Future trends In the past few decades, the dairy industry has moved from being solely a commodity food-based industry producing significant waste (whey) volumes to one that has increasingly sought added value and the minimization of waste. The nature of dairy products has expanded from simple nutritional © Woodhead Publishing Limited, 2010
474 Separation, extraction and concentration processes foods to separate fractions sold as individual ingredients or as blends to meet the specific functional and nutritional requirements of secondary producers. This trend will continue, as milk volumes increase to meet the new demand for dairy products from the developing world and greater bulk production efficiencies drive companies to find new products to maintain their competitive advantages. Nutraceutical products, with implicit claims of health benefits, are now commonplace and there are likely to be more such products in the future, with some fractions in the future perhaps even meeting the more stringent proof of efficacy requirements of pharmaceuticals for highly targeted applications, for example around neonatal gut development, anticancer agents or immune system enhancers. The commercial recovery of very minor fractions with useful bioactivities, such as oligosaccharides (including those incorporating sialic acid), the myriad proteins that make up the milk fat globule membrane, and the complex lipids therein (phosphatidylcholine, phosphatidylethanolamine, phosphatidylserine, phosphatidylinositol, sphingomyelin and glycolipids such as gangliosides) will be a challenge. Recovery of such fractions may require specific affinity techniques and the challenge will be to use techniques to extract minor proteins economically on a large scale, while maintaining the safety and integrity of the residual milk or whey for further processing i.e. without contaminating the flow-through stream with leachates. On-farm extraction of minor products from milk may offer advantages in the production of proteins and other bioactives that are heat-labile and are currently largely lost during the pasteurization required inside the factory gate. On-farm extraction may also offer direct benefit to farmers who select for certain genetic traits or apply unique feeding regimes to enhance the production of specific high-value products in their herds. In summary, higher volumes of milk production and increased global competition in the dairy industry will probably drive further growth in niche bioactive products over the next few decades. High-volume, economically efficient separation technologies with high specificity and compatibility with food production will be required to support this growth.
16.9 References Abraham, G. B. (1988). Process for preparing antibodies against E. coli K-99 antigen from bovine milk, US patent 4784850. Akita, E. M. and Li-Chan, E. C. Y. (1998). ‘Isolation of bovine immunoglobulin G subclasses from milk, colostrum, and whey using immobilized egg yolk antibodies’. Journal of Dairy Science 81(1): 54–63. Al-Mashikhi, S. A., Li-Chan, E. and Nakai, S. (1988). ‘Separation of immunoglobulins and lactoferrin from cheese whey by chelating chromatography’. Journal of Dairy Science 71(7): 1747–1755. Al-Mashikhi, S. A. and Nakai, S. (1987). ‘Isolation of bovine immunoglobulins and lactoferrin from whey proteins by gel filtration techniques’. Journal of Dairy Science 70(12): 2486–2492.
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480 Separation, extraction and concentration processes Brazilian Journal of Medical and Biological Research 19(2): 149–157. Saleh, Z. S. and Hossain, M. M. (2001). ‘A study of the separation of proteins from multicomponent mixtures by a semi-batch foaming process’. Chemical Engineering and Processing 40(4): 371–378. Saufi, S. M. and Fee, C. J. (2009). ‘Fractionation of b-lactoglobulin from whey by mixed matrix membrane ion exchange chromatography’. Biotechnology and Bioengineering 103(1): 138–147. Sauter, E. A. and Montoure, J. E. (1972). ‘The relationship of lysozyme content of egg white to volume and stability of foams’. Journal of Food Science 37(6): 918–920. Schlatterer, B., Baeker, R. and Schlatterer, K. (2004). ‘Improved purification of b-lactoglobulin from acid whey by means of ceramic hydroxyapatite chromatography with sodium fluoride as a displacer’. Journal of Chromatography B: Analytical Technologies in the Biomedical and Life Sciences 807(2): 223–228. Schmerr, M. J. F., Patterson, J. M., Van Der Maaten, M. J. and Miller, J. M. (1985). ‘Conditions for binding bovine IgG1 to protein A-sepharose’. Molecular Immunology 22(5): 613–616. Seifu, E., Buys, E. M. and Donkin, E. F. (2005). ‘Significance of the lactoperoxidase system in the dairy industry and its potential applications: A review’. Trends in Food Science and Technology 16(4): 137–154. Semo, E., Kesselman, E., Danino, D. and Livney, Y. D. (2006). Casein micelle as a natural nano-capsular vehicle for nutraceuticals. Food Colloids 2006, April 23–26, Montreux, Switzerland. Severin, S. and Xia, W. S. (2005). ‘Milk biologically active components as nutraceuticals: Review’. Critical Reviews in Food Science and Nutrition 45(7–8): 645–656. Shea, A. P., Crofcheck, C. L., Payne, F. A. and Xiong, Y. L. (2009). ‘Foam fractionation of a-lactalbumin and b-lactoglobulin from a whey solution’. Asia-Pacific Journal of Chemical Engineering 4(2): 191–203. Shin, K., Hayasawa, H. and Lonnerdal, B. (2001). ‘Purification and quantification of lactoperoxidase in human milk with use of immunoadsorbents with antibodies against recombinant human lactoperoxidase’. American Journal of Clinical Nutrition 73(5): 984–989. Shiozawa, M., Okabe, H., Nakagawa, Y., Morita, H. and Uchida, T. (2001). ‘Purification of lactoferrin by expanded-bed column chromatography’. Kagaku Kogaku Ronbunshu 27(2): 147–148. Spitsberg, V. L. (2005). ‘Bovine milk fat globule membrane as a potential nutraceutical’. Journal of Dairy Science 88(7): 2289–2294. Splitt, H., Mackenstedt, I. and Freitag, R. (1996). ‘Preparative membrane adsorber chromatography for the isolation of cow milk components’. Journal of Chromatography A 729(1–2): 87–97. Stec, J., Bicka, L. and Kuźmak, J. (2004). ‘Isolation and purification of polyclonal IgG antibodies from bovine serum by high performance liquid chromatography’. Bulletin of the Veterinary Institute in Pulawy 48(3): 321–327. Su, C. K. and Chiang, B. H. (2003). ‘Extraction of immunoglobulin-G from colostral whey by reverse micelles’. Journal of Dairy Science 86(5): 1639–1645. Suutari, T. J., Valkonen, K. H., Karttunen, T. J., Ehn, B. M., Ekstrand, B., Bengtsson, U., Virtanen, V., Nieminen, M. and Kokkonen, J. (2006). ‘IgE cross reactivity between reindeer and bovine milk b-lactoglobulins in cow’s milk allergic patients’. Journal of Investigational Allergology and Clinical Immunology 16(5): 296–302. Syväoja, E.-L., Ahola-Luttila, H. K., Kalsta, H., Matilainen, M. H., Laakso, S., Husu, J. R. and Kosunen, T. U. (1994). ‘Concentration of campylobacter-specific antibodies in the colostrum of immunized cows’. Milchwissenschaft 49: 27–31. Tellez, C. M. and Cole, K. D. (2000). ‘Preparative electrochromatography of proteins in various types of porous media’. Electrophoresis 21(5): 1001–1009. Thomassen, E. A. J., Van Veen, H. A., Van Berkel, P. H. C., Nuijens, J. H. and Abrahams,
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482 Separation, extraction and concentration processes Yang, H., Gurgel, P. V. and Carbonell, R. G. (2005). ‘Hexamer peptide affinity resins that bind the Fc region of human immunoglobulin G’. Journal of Peptide Research 66(Suppl. 1): 120–137. Ye, X., Yoshida, S. and Ng, T. B. (2000). ‘Isolation of lactoperoxidase, lactoferrin, a-lactalbumin, b-lactoglobulin B and b-lactoglobulin A from bovine rennet whey using ion exchange chromatography’. International Journal of Biochemistry and Cell Biology 32(11–12): 1143–1150. Yoshida, S. (1989). ‘Preparation of lactoferrin by hydrophobic interaction chromatography from milk acid whey’. J. Dairy Sci. 72: 1446–1450. Zydney, A. L. (1998). ‘Protein separations using membrane filtration: new opportunities for whey fractionation’. International Dairy Journal 8(3): 243–250.
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17 Methods of concentration and purification of omega-3 fatty acids S. P. J. Namal Senanayake, Danisco USA, Inc., USA
Abstract: An overview is presented of the various methodologies used for producing highly purified omega-3 fatty acids from natural source materials. Omega-3 fatty acids derived from fish, krill and microalgae, consisting of eicosapentaenoic acid (EPA; 20:5n-3), and docosahexaenoic acid (DHA; 22:6n-3), have beneficial effects in the prevention and management of cardiovascular disease and other chronic disorders. Production of high-purity omega-3 fatty acids is increasingly important in both the nutraceutical and pharmaceutical industries. The physical, chemical and enzymatic methods used include urea adduction, chromatography, low-temperature fractional crystallization, supercritical fluid extraction and distillation. Key words: omega-3 fatty acids, urea adduction, chromatography, crystallization, supercritical fluid extraction, distillation, enzymatic methods.
17.1 Introduction Omega-3 fatty acids are essential fatty acids that have diverse biological effects in human health and disease. They are essential to human health but cannot be manufactured by the body. For this reason, omega-3 fatty acids must be obtained from food. Omega-3 polyunsaturated fatty acids (PUFAs) can be found in fish, such as salmon, tuna, and halibut, other marine organisms such as algae and krill, certain plants, and nut oils. The heart-health benefits of the omega-3 fatty acids are well known. Omega-3 fatty acids are considered to have beneficial effects in the prevention of cardiovascular disease, inflammation, hypertension as well as other chronic disorders (Kris-Etherton et al., 2001; Lands, 2003; Senanayake, 2000; Senanayake and Shahidi, 2000b; Shahidi and Senanayake, 2006; Yokoyama et al. 2007). In 2004, the US Food and © Woodhead Publishing Limited, 2010
484 Separation, extraction and concentration processes Drug Administration gave ‘qualified health claim’ status to eicosapentaenoic acid (EPA; 20:5n-3) and docosahexaenoic acid (DHA; 22:6n-3) found in fish, stating that ‘supportive but not conclusive research shows that consumption of EPA and DHA omega-3 fatty acids may reduce the risk of coronary heart disease’ (United States Food and Drug Administration, 2004). For clinical and nutritional applications, the natural sources of omega-3 fatty acids, as such, may not provide the required amounts of these fatty acids and hence concentration and purification of omega-3 fatty acids may be necessary. A concentrated source of these fatty acids is desired to achieve sustainable benefits. Highly purified omega-3 fatty acids may be produced in the free fatty acid, simple alkyl ester and triacylglycerol forms. To achieve this, physical, chemical and enzymatic techniques may be employed. Methods traditionally employed for the concentration and purification of omega-3 fatty acids include urea adduction, chromatographic methods, low-temperature fractional crystallization, supercritical fluid extraction, distillation, and enzymatic and integrated methods.
17.2 Urea adduction in the concentration and purification of omega-3 fatty acids Urea adduction is one of the most efficient and simplest techniques for concentration and purification of omega-3 fatty acids from natural sources (Senanayake, 2000). The formation of complexes between urea and straightchain saturated fatty acids is a well established and potentially valuable separation technique for fractionation of free fatty acids or esters (Shahidi and Senanayake, 2006). Initially, the triacylglycerols (TAGs) of oil are hydrolyzed into their constituent fatty acids via alkaline hydrolysis using alcoholic KOH or NaOH. The resultant free fatty acids (FFAs) are then mixed with an ethanolic solution of urea for complex formation. Urea molecules readily form solid-phase complexes with saturated and monounsaturated fatty acids and crystallize out on cooling and may be removed by filtration. The liquid fraction is highly enriched with omega-3 fatty acids. In general, the crystallization temperature can be ranged from ambient to –20 °C. Since the process involves relatively mild operating conditions and chemicals (free fatty acids, urea, ethanol, water) that are generally recognized as safe (GRAS) by the US Food and Drug Administration, it can be labeled as eco-friendly. This process can be used as a preliminary fractionation step in conjunction with other purification methods, such as low-temperature fractional crystallization, enzyme-catalyzed methods, and molecular distillation, resulting in a highpurity free-fatty-acid omega-3 product. Urea-based fractionation of fatty acids has been extensively used for enriching omega-3 fatty acids in marine oils (Hayes et al., 2000). Urea adduction of fatty acid ethyl esters of squid (Illex argentinus) visceral oil,
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using cyclohexane and methanol, can increase the EPA content from 11.8 to 28.2% and DHA content from 14.9 to 35.6% (Hwang and Liang, 2001). Haagsma et al. (1982) described a urea adduction method for enriching the EPA and DHA levels of cod liver oil from 12 to 28 and 11 to 45%, respectively. Wanasundara and Shahidi (1999) used an experimental design to optimize the conditions on a laboratory scale that led to the maximum concentration of EPA and DHA from seal blubber oil, using the urea-based fractionation. The authors obtained 88.2% of total omega-3 fatty acids at an urea/fatty acid ratio of 4.5, a crystallization time of 24 h and a crystallization temperature of –10 °C. Senanayake and Shahidi (2000a) concentrated and purified DHA from the oil extracted from the microalgae Crypthecodinium cohnii and reported a DHA enrichment from 47.4 to 97.1% with a process yield of 32.5% of the mass of the original algal oil. The fatty acid composition of algal oil and its DHA concentrate obtained by urea adduction are reported in Table 17.1 (Senanayake and Shahidi, 2000a). In another study, mackerel processing waste comprising skins, viscera, and muscle tissue was evaluated by Zuta et al. (2003) for concentrating omega-3 fatty acids by urea adduction. Fish oil was extracted using either chloroform/methanol (2:1 vol/vol) or hexane/isopropanol (3:2 vol/vol). The oil yield and iodine value (which measures the degree of unsaturation in fats and oils) were determined for fresh fish oil extracts. Omega-3 fatty acid concentrates were prepared from saponified fish oil via urea adduction. The mean oil yields were 9.18, 9.2, and 38.1% for viscera, muscle, and skin, respectively. The mean baseline iodine value was 134, which increased to 296 after urea adduction. Hence, it was possible to concentrate omega-3 fatty acids from mackerel processing waste. The type of tissue used did not affect the amount of omega-3 fatty acids concentrated. Mackerel skin was most desirable because of its high oil content. Hayes (2006) investigated the effect of cooling rate on the degree of removal of saturated acyl groups from FFAs derived from canola oil and Table 17.1 Fatty acid composition of algal oil and its DHA concentrate after urea adduction (adapted from Senanayake and Shahidi, 2000a) Fatty acids 10:0 12:0 14:0 16:0 16:1 18:1w-9 18:2 w-6 22:5 w-3 22:6 w-3 Iodine value, calculated
Algal oil (%) 0.6 1.1 15.0 9.0 2.2 19.0 1.0 0.5 47.4 234
DHA concentrate (%) 0.5 0.5 0.1 0.0 0.3 0.2 0.7 0.4 97.1 437
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486 Separation, extraction and concentration processes the isolation of di- and polyunsaturated acyl groups from FFAs derived from vegetable and fish oil during urea adduction. Traditionally, slow cooling has been used (–1 °C min–1). A more rapid cooling rate (–47 °C min–1) produced urea inclusion compounds (crystals) of similar morphology and thermodynamic properties, but of a size an order of magnitude smaller than the urea inclusion compounds formed during slow cooling. Fractionations used only renewable materials (urea, FFAs, and 95% ethanol as solvent) and benign operating conditions (ambient pressure, 25–75 °C, and neutral pH). When the recovery of FFAs was relatively high (>60%), the selectivity of urea adduction toward the inclusion of saturated fatty acids and against PUFAs was not affected by the cooling rate. In contrast, when the FFAs recovery was low, representing those instances where an increase in the purity of the PUFAs is a more important economic goal, a slower cooling rate resulted in a significantly greater discrimination against PUFA groups, hence to a FFA product with a measurably greater purity.
17.3 Chromatographic methods for the concentration and purification of omega-3 fatty acids Another method for concentration and purification of omega-3 fatty acids is the use of chromatography. High-performance liquid chromatography, silver resin chromatography and supercritical-fluid chromatography have been used for concentration of omega-3 fatty acids from natural sources. Hayashi and Kishimura (1993) isolated 63–74% pure DHA from skipjack tuna eye orbital oil by stepwise elution with hexane, diethyl ether/hexane and diethyl ether on a silicic acid column. Teshima et al. (1978) employed a silver nitrate-impregnated silica gel column to separate DHA and EPA from squid-liver oil after forming methyl esters. They were able to isolate 95–98% DHA and 85–96% EPA with yields of 48 and 39%, respectively. Guil-Guerrero and Belarbi (2001) purified EPA and DHA from cod liver oil using a silver nitrate-impregnated silica gel column. The oil was saponified and treated with urea, and the non-complexed fatty acids were then converted to methyl esters before chromatography. The column was washed with a sequence of mobile phases. A 64% recovery of DHA with 100% purity was obtained. The recovery of EPA was 29.6%, with a final purity of 90.6%. Perretti et al. (2007) investigated the fractionation of fish oil fatty acid ethyl esters with the aim of obtaining a lipid fraction enriched in omega-3 fatty acids and with a suitable EPA/DHA ratio. They reported the possibility of modifying the original fatty acid ethyl ester concentrations by optimizing the extraction conditions in terms of pressure, temperature, and supercritical CO2 flow rate: 2-h runs, pressures of 100, 140, 150, and 300 bar, and liquid CO2 flow rates of 2.5, 3.5, 5, and 10 kg h–1, at 40, 50, and 60 °C, in the three sections of the column starting from the bottom, respectively. They
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stated that supercritical fluid fractionation appears to be a useful processing technique for changing the composition of lipids in order to obtain high added-value functional products. A preparative reversed-phase high-performance liquid chromatography method with gradient elution using acetonitrile–chloroform and evaporative light-scattering detection was used by Mansour (2005) to purify milligram quantities of microalgal PUFAs, separated as methyl esters. PUFA-methyl esters purified included methyl esters of DHA, EPA and the unusual very-longchain highly unsaturated fatty acid, octacosaoctaenoic acid (28:8n-3) from the marine dinoflagellate, Scrippsiella sp. Other fatty acids purified from various microalgae using this method to greater than 95% purity included 16:3(n-4), 16:4(n-3), 16:4(n-1) and 18:5(n-3). The number of injections required was variable and depended on the abundance of the desired PUFA-methyl esters, and resolution from closely eluting PUFA-methyl esters, which determined the maximum loading. The purity of these fatty acids was determined by electron-impact gas chromatography–mass spectrometry (EI GC–MS) and the chain length and location of double bonds were determined by EI GC– MS of 4,4-dimethyloxazoline derivatives formed using a low-temperature method. The advantages over silver-ion HPLC for purifying PUFA-methyl esters are that separation occurs according to chain length as well as degree of unsaturation enabling separation of PUFA-methyl esters with the same degree of unsaturation but different chain length. In addition, PUFA-methyl esters were not strongly adsorbed, but eluted earlier than their more saturated corresponding PUFA-methyl esters of the same chain length. This method is robust, simple, and requires only a short re-equilibration time. Centrifugal partition chromatography (CPC) is a separation method based on the liquid partition of compounds. Owing to the high centrifugal field force, one phase stays in the rotor (the liquid stationary phase), the other one is the mobile phase as in classical liquid chromatography. CPC has been used for isolation and purification of PUFAs from various natural sources. Several important advantages of centrifugal partition chromatography are reported in Table 17.2. A CPC method to purify DHA from microalgal oil on a laboratory scale, has been developed by Wanasundara and Fedec (2002). The starting microalgal oil contained 39.7% DHA and 15.2% DPA (n-6), along with several other fatty acids. The mobile phase was hexane–methanol–water in a normal phase-ascending mode. A good separation was achieved and recoveries of 84.6% DHA and 84.9% DPA were obtained. Bousquet and Goffic (1995) have examined CPC separation of microalgal oil-based DHA and EPA and were able to isolate pure DHA and EPA from this oil with excellent yields. The first separation used heptane as the stationary phase and aqueous 3% acetonitrile as the mobile phase. The minor fatty acids were eluted, leaving a mixture of four major PUFAs. This mixture was subjected to a second separation, using heptane as a stationary phase and aqueous methanol as the mobile phase. Under these conditions, it was possible to isolate pure DHA and EPA with high yields. High-speed countercurrent chromatography was used by Du et
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488 Separation, extraction and concentration processes Table 17.2 Advantages of centrifugal partition chromatography Advantages Cost
Solid supports, which in many cases are costly, are not required but instead a solvent system is used Recovery Elimination of solid supports avoids problems associated with irreversible retention of highly retentive sample components. Almost 100% recovery of the compounds is guaranteed Quantity Compared with conventional liquid chromatography, the volume ratio of the stationary phase to the total column (rotor) volume is greater. Hence, large quantities of materials can be retained in the stationary phase Speed The stationary phase is retained by centrifugal force, enabling the mobile phase to be pumped at high speeds through the apparatus resulting in reduced separation times Stability Mild operating conditions are used. Therefore, decomposition and oxidation of PUFAs are virtually nonexistent under these conditions Versatility Any two-phase solvent system can be used, including many prepared from nontoxic and commonly available solvents
al. (1996) to separate DHA and EPA ethyl esters in a mixture containing 39.3 and 56.4%, respectively, with hexane–dichloromethane–acetonitrile (5:1:4 v/v/v) as mobile phase. However, the separation was not successful with this solvent system because of the low partition coefficients of the esters under the experimental conditions used. Murayama et al. (1988) separated a mixture of stearic, oleic, linoleic and a-linolenic (18:3n-3) acid ethyl esters because their partition coefficients were distributed over a wide range in the mobile phase comprising hexane–acetonitrile (1:1 v/v). The ethyl esters of linoleic acid and a-linolenic acids were successfully separated during the first normal ascending elution whereas the ethyl ester of oleic acid was isolated by switching the elution mode to descending. Despite the developments in chromatographic techniques to concentrate and purify omega-3 oils, the use of very large volumes of solvents, loss of column resolution after repeated use, and potential product solvent residues are likely to hinder scale-up to production scale volumes.
17.4 Low-temperature fractional crystallization for the concentration and purification of omega-3 fatty acids Low-temperature fractional crystallization is one of the simplest methods employed for production of purified omega-3 fatty acids. This process takes advantage of the existing differences in the melting points of different fatty acids as neat oils or in different solvent systems. The more saturated fatty acids have higher melting points and crystallize out of the mixtures
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leaving behind the more unsaturated fatty acids (Senanayake, 2000). The melting points of fatty acids are dependent on their degree of unsaturation. For example, EPA and DHA melt at –54 and –44.5 °C compared with 13.4 and 69.6 °C for 18:1 and 18:0, respectively (Merck Index, 1983). As the temperature of a mixture of saturated and unsaturated fatty acids decreases, the saturated fatty acids, having a higher melting point, start to crystallize out first and the liquid phase becomes enriched in the unsaturated fatty acids (Shahidi and Senanayake, 2006). However, as the number and type of fatty acid components in the mixture increases, the crystallization process becomes more complex and repeated crystallization and separation of fractions must be carried out to obtain purified fractions. For fish oils, not only is there a very wide spectrum of fatty acids but the fatty acids exist, not in the FFA form, but esterified in TAGs. However, the principle of low-temperature fractional crystallization can still be applied to marine oils to partially concentrate TAGs rich in omega-3 PUFAs (Shahidi and Wanasundara, 1998). Seal blubber oil in the TAG and FFA forms were subjected to low-temperature fractional crystallization using solvents such as hexane and acetone in order to obtain omega-3 fatty acid concentrates. Preparation of omega-3 fatty acid concentrates from seal blubber oil by low-temperature crystallization was also reported by Wanasundara (1996), who subjected seal blubber oil in the TAG or FFA form to solvent-based fractionation, using hexane and acetone as solvents, at different temperatures. The content of omega-3 fatty acids in the non-crystalline fraction was increased with decreasing the crystallization temperature. Under all temperature conditions, acetone gave rise to the highest concentration of total omega-3 fatty acids. Low-temperature solvent crystallization of seal blubber oil, in the FFA form, at –60 and –70 °C in hexane, resulted in total omega-3 fatty acid content of up to 58.3 and 66.7%, respectively. However, the content of total omega-3 fatty acids in acetone increased up to 56.7 and 46.8%, respectively. In another study, Han et al. (1987) found that alkali salts of less unsaturated fatty acids crystallize more rapidly than those of highly unsaturated fatty acids containing four or more double bonds, when the saponified solution is cooled. They also compared the cooling temperature and the rate of cooling on the enrichment of omega-3 fatty acids of alkali salts of sardine oil fatty acids. The DHA and EPA from sardine oil were concentrated more than 2.3fold with minimum yields of 91 and 87%, respectively. Fatty acid profiles of the prepared concentrates showed that the cooling rate and temperature had little effect on the yield and contents of DHA and EPA. Chen and Ju (2001) utilized a modified low-temperature solvent crystallization process for the enrichment of PUFAs in borage and linseed oil fatty acids. The effects of solvent, operation temperature, and solvent-to-FFA ratio on the concentration of PUFAs were investigated. The best results were achieved when a mixture of 30% acetonitrile and 70% acetone was used as the solvent. With an operating temperature of −80 °C and a solvent-to-FFA ratio of 30 mL g –1, g-linolenic acid (GLA; 18:3n-6) in FFA of saponified borage oil was raised
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490 Separation, extraction and concentration processes from 23.4 to 88.9% with a yield of 62.0%. At a yield of 24.9%, a-linolenic acid (ALA; 18:3n-3) in linseed oil was increased from 55.0 to 85.7%.
17.5 Supercritical-fluid extraction for the concentration and purification of omega-3 fatty acids Supercritical-fluid extraction is a relatively novel technique which has found use in food and pharmaceutical applications (Senanayake, 2000). The properties of a supercritical fluid are used to selectively extract a specific compound or to fractionate mixtures by changing the temperature and pressure without any phase change. The most important property of these fluids is the high solvation capacity in their supercritical region (Shahidi and Senanayake, 2006; Shahidi and Wanasundara, 1998). This method is mild and, because it uses CO2, minimizes autoxidation (Senanayake, 2000). A number of gases are known to have good solvent properties at pressures above their critical values. For food applications, CO2 is the solvent of choice because it is inert, inexpensive, non-flammable, environmentally acceptable, non-toxic, relatively safe, completely recoverable, and readily available and has a moderate critical temperature (31.1 °C) and pressure (1070 psig). The fatty acids were most effectively separated on the basis of chain length; hence the method works best for oils with low levels of long-chain fatty acids. Because this method is derived from separation of compounds based on their molecular weight and not their degree of unsaturation, prior concentration steps, such as urea adduction or low-temperature crystallization, may be necessary in order to purify the omega-3 fatty acids. The use of supercritical fluids for extraction and purification of omega-3 fatty acids from fish and krill oils has been reported (Mishira et al., 1993; Yamagouchi, et al., 1986). Fish oils in the form of free fatty acids and fatty acid esters have been extracted with supercritical CO2 to yield concentrates of EPA and DHA. The use of high pressures and high capital costs might limit the widespread use of this technique in large-scale applications (Shahidi and Senanayake, 2006). Alkio et al. (2000) evaluated the technical and economic feasibility of producing ethyl ester concentrates of DHA and EPA from transesterified tuna oil using supercritical fluid chromatography. A systematic experimental procedure was used to find the optimal values for process parameters and the maximal production rate. DHA ester concentrates up to 95 wt% purity were obtained in one chromatographic step with supercritical fluid chromatography, using CO2 as the mobile phase at 65 °C and 145 bar and octadecylsilane-type reversed-phase silica as the stationary phase. DHA ester, 0.85 g/(kg stationary phase/h) and 0.23 g EPA ester/(kg stationary phase/h) can be simultaneously produced at the respective purities of 90 and 50 wt%. The process for producing 1000 kg of DHA concentrate and 410 kg of EPA concentrate per year requires 160 kg of stationary phase
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Methods of concentration and purification of omega-3 fatty acids 491
and 2.6 t h–1 of CO2 as recycling mobile phase. The operating cost of the supercritical fluid chromatography was US$550 kg–1 for DHA and EPA ethyl ester concentrate. The US Patent 4,675,132 (Stout and Spinelli, 1987) demonstrated that fish oil esters could be fractionated by supercritical fluid extraction to produce an oil with a DHA content of 60–65%. However, the recovery of omega-3 fatty acids was low during this process. Jachmanián et al. (2007) studied the solubility of various ethyl esters derived from hake liver oil in supercritical CO2. A selectivity factor was used to determine optimal conditions to fractionate the ethyl ester mixture. A strong influence of solvent pressure and temperature was observed at 8.63–18.04 MPa and 40–70 °C. The lowest total solubility of the ethyl ester mixture was obtained when using supercritical CO2 at the lowest density (the lowest pressure and the highest temperature value tested). The highest discrimination against longchain PUFAs (e.g. EPA and DHA) was also obtained under these conditions. Conversely, higher solubility and lower selectivity were obtained when solvent density increased. Considering this inverse correlation between selectivity and solubility, a single-step batch-fractionation process was designed to increase the DHA ethyl ester content from an initial value of 17.5% in the starting material to 55% in the final extract. Davarnejad et al. (2008) examined the solubility of fish oil in supercritical CO2 at temperatures of 40, 50, 60, and 70 °C and pressures of 13.6, 20.4, and 27.2 MPa. The fractionated fish oil samples collected were then esterified using methanol with sodium methoxide catalyst. The samples were analyzed by GC to determine the amount of four fatty acid methyl ester (FAME) components extracted, namely, methyl palmitate, methyl oleate, methyl EPA (5,8,11,14,17-eicosapentanoate), and methyl DHA (4,7,10,13,16,19docosahexenoate). The results showed that the highest solubility of the fish oil (0.921 g of oil in 100 g of CO2) was obtained at optimum conditions of 40 °C and 27.2 MPa. The solubility of fish oil in supercritical CO2 was found to be higher at lower temperature and at lower fractionation time. Furthermore, the average yield obtained for the combined total of the four FAME components was 66%, with methyl palmitate having the highest at 30.5% under extraction conditions of 50 °C and 13.6 MPa whereas methyl EPA has the lowest at 3.24%. Létisse et al. (2006), and Létisse and Comeau (2008) evaluated the enrichment of EPA and DHA from sardine heads, a waste product from the fish canning industry, via supercritical fluid extraction. These studies were done on a laboratory scale. Various parameters, such as pressure, temperature, CO2 rate and time were optimized in order to obtain the highest yield of extracted oil with the highest amount of EPA and DHA in the extraction product. In the first approach, the oil yield was measured. Then, a quadratic model with three variables was employed to maximize the EPA and the DHA concentrations. A multicriteria optimization, using the desirability function, was performed to determine the best level for each parameter. Pressure, temperature and CO2 rate were, respectively, set at 300 bar, 75 °C and © Woodhead Publishing Limited, 2010
492 Separation, extraction and concentration processes 2.5 ml min–1 during the 45 min extraction. A yield of 10.36% of extracted oil (compared with the dry material) was achieved with an amount of 10.9% of EPA and 13.0% of DHA (compared with all fatty acids of the extract). These yields were lower than with a traditional solvent extraction. However, the advantages of supercritical fluid extraction were shorter extraction time, prevention of heating, and better organoleptic properties by excluding the use of toxic organic solvents. Catchpole et al. (2000) reported the countercurrent extraction and fractionation of a range of crude fish oils using supercritical CO2 and CO2– ethanol mixtures. Vitamin A palmitate was extracted from model mixtures of cod liver oil and vitamins using pure CO2. The separation factor was low, owing to similar solubilities of the vitamin ester and the oil. Vitamin A was also recovered from cod liver oil ethyl esters–vitamin A mixtures. The separation factor was substantially improved over the non-esterified oil, owing to large differences in the solubilities of the esters and vitamin A in supercritical CO2. Solubilities of fish oils and squalene are reported using CO2–ethanol mixtures at 333 K, ethanol concentrations of 0 to 12% by mass, and pressures of 200–300 bar. Solubilities of all oils and squalene increased exponentially with linear increases in the ethanol concentration. The solubility of polar components increased more rapidly than non-polar components. Pilot-scale removal of fatty acids from Orange Roughy oil and squalene from deep sea shark liver oil was carried out using CO2–ethanol mixture as the solvent. The extent of fatty acid removal from Orange Roughy oil was higher than with pure CO2, whereas the degree of separation of squalene from shark liver oil was lower. However, throughput was substantially increased relative to pure CO2 in both cases. Temelli et al. (1995) optimized the supercritical CO2 extraction temperature and pressure for oil removal from freeze-dried, fall Atlantic mackerel. The effect of extraction conditions on pH and water-binding potential of the protein residue was evaluated. For the temperature range (35–55 °C) and pressure interval (20.7–34.5 MPa), a combination of 34.5 MPa and 35 °C gave the highest oil yield and concentration of omega-3 fatty acids.
17.6 Distillation methods for the concentration and purification of omega-3 fatty acids Fractional distillation is another simple process for separation of mixtures of omega-3 fatty acids under ultralow pressure. This method takes advantage of differences in the boiling point and molecular weight of fatty acids under the conditions of high temperature (180–250 °C) and reduced pressure (0.1 to 1.00 mm Hg). Fatty acid enrichment is achieved by exploiting the differences in vapour pressure through countercurrent contacting of vapour
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and liquid phases in stages using plates or continuously using random or structured packing. Molecular distillation, often also called ‘short path distillation’, may use moderately lower temperatures and short heating intervals. Molecular distillation provides high selectivity, but as a result of its relatively high operating temperatures and reduced pressures, it leads to high operating costs, possible thermal degradation of omega-3 fatty acids and safety concerns. Several advantages and disadvantages of molecular distillation are reported in Table 17.3. Molecular distillation can be used for the concentration of omega-3 fatty acids from fish oils. The content of omega-3 fatty acids in fish oils mainly depends on the fish species itself, their habitat and the season. A typical base material for commercial production of omega-3 fatty acid concentrates is anchovy or sardine oil with a content of approximately 18% EPA and 12% DHA. Tuna oil is also used for the large-scale production of DHA concentrates, as the crude oil contains up to 25% DHA. In general, the concentration of omega-3 fatty acids in the crude oil should not be less than 30% in total. The state-of-the-art method for the concentration of omega-3 fatty acids from fish oil by molecular distillation is by the fractionation of their ethyl esters (Fig. 17.1). The latter can be obtained by interesterification of fish oil with ethyl alcohol. The equipment used in this process consists of two stills (two-stage molecular distillation) with a degasser. The product distillation requires a minimum of three passes through the system. The first, degassing pass, removes any moisture left after interesterification. The residue from this pass is then used as the feed for the second, ‘lights’ removal, pass. The second pass concentrates the DHA and EPA by separating and removing the ‘lights’ from the feed. The second pass removes 20–50% of C10–C18 fatty acid esters. The residue from this pass is then used as the feed for the third, ‘product distillation’, pass. The third pass further concentrates the DHA and EPA (up to 40–80%) by separating the heavier fraction from the feedstock. The heavy fraction (5–10%) consists of longer chain fatty acids, including fatty acids chain length >C24 as residue. Operating pressures and temperatures required are 0.005–0.01 torr and 170–190 °C, respectively. The recoveries can be up to 70% at DHA and EPA concentrations of 55–65 wt%. Table 17.3 Advantages and disadvantages of molecular distillation Advantages
Disadvantages
Stability: the vacuum allows oils to be processed at minimal temperatures, reducing the risk of thermal decomposition and oxidation of PUFAs
Cost: cost is relatively high
Purity: separating the oil components by weight allows contaminants to be reduced far below industry specifications
Natural form: the starting natural triglyceride form is lost in the process
Concentration: weight grouping allows the processor to concentrate specific fatty acids
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494 Separation, extraction and concentration processes Crude fish oil
Refining, bleaching and deodorization
Refined fish oil
Interesterification
Fish oil ethyl esters
Molecular distillation
Purified omega-3 fatty acid esters
Fig. 17.1 Purification of crude fish oil into highly concentrated omega-3 fish oil esters.
The DHA to EPA ratio primarily depends on their content in the base material and on the degree of concentration, as EPA ethyl ester, the more volatile component, also accumulates in the mid-chain-length fatty acid ester fraction. Reprocessing of the latter fraction by molecular distillation increases the recovery of omega-3 fatty acids. Stout et al. (1990) pointed out the practical difficulty of purifying omega-3 fatty acids from menhaden oil via molecular distillation in the natural TAG form. The distillation of menhaden oil, in its natural TAG form, concentrated only EPA from an initial value of 16.0% to 19.5%. Hence, this study did not result in a significant improvement in the concentration of omega-3 fatty acids, as they are more or less uniformly distributed in the TAGs. Furthermore, distillation of triglycerides require high temperatures, risking thermal decomposition of unsaturated fatty acids, and a very low operating pressure, which results in an inefficiently low specific feed rate. However, molecular distillation of its ethyl esters increased the EPA content from 15.9% to 28.4% (Stout et al., 1990). The degree of concentration of DHA was even more remarkable. Whereas DHA doubled from 8.4% to 17.3% in the TAG form, in the alkyl ester form it increased from 9.0% to 43.9% (Stout et al. 1990). Although the greater volatility of the fatty acid esters allows the use of lower temperatures (compared with process temperatures required for FFA fractionation), temperatures are still moderately high, and exposure to distillation conditions over a prolonged period of time can be © Woodhead Publishing Limited, 2010
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detrimental to the omega-3 fatty acid constituents, causing polymerization, hydrolysis, isomerization and thermal degradation. Methyl esters of Atlantic herring oil containing 62% monoethylenic fatty acids were subjected to batch fractional distillation, under vacuum on a pilot-plant scale, to study the feasibility of fractionating fatty acid esters of marine oils of low iodine value into monounsaturated fractions with increased commercial value for industrial chemical uses (Ackman et al., 1973). A total of 64 methyl ester fractions were collected and recoveries of the major saturated and monounsaturated acids were close to 100%, and some fractions contained over 90% of the desired 22:1 long chain monounsaturated acids. The short-chain polyunsaturated acids were recovered in high yields, but the long-chain highly unsaturated acids were recovered in yields of 60% or less owing to oxidative and thermal decomposition in the particular apparatus employed. If small amounts of unsaturated acids are acceptable, fractional distillation of low iodine value marine oils could inexpensively supply fractions with high concentrations of methyl esters of longer chain (C20 and C22) monounsaturated and shorter chain (C14) saturated acid or (C16) saturated–monounsaturated acid mixture. Liang and Hwang (2000) employed short-path distillation to fractionate EPA and DHA ethyl esters from squid visceral oil. The elimination temperatures of squid visceral oil ethyl esters (SVOEE) ranged from 50 to 140 °C, increasing with the carbon number of ethyl esters. The elimination temperature of cholesterol was higher than those of SVOEE. When SVOEE originally containing 9.0% EPA, 14.7% DHA, and 11.21 mg g–1 of cholesterol was distilled from 50 to 150 °C, the 130 °C distillate contained 15.5% EPA and 34.7% DHA with 0.99 mg g–1 of cholesterol, and the yield was 21.8%. The 150 °C distillate had 43.1% DHA with 4.96 mg g–1 of cholesterol. Furthermore, the distillates collected from 110 to 150 °C contained 24.4 to 50.2% of EPA plus DHA, and their total yield was 58.3%.
17.7 Enzymatic methods for the concentration and purification of omega-3 fatty acids For the concentration of omega-3 fatty acids on a large scale, each of the above physical and chemical separation methods have some drawbacks either in terms of low yield, a requirement for large volumes of solvent or sophisticated equipment, a risk of structural changes in the fatty acid products, or high operational costs (Senanayake, 2000). Lipases work under mild conditions of temperature and pH (Gandhi, 1997), a factor which favors their potential use for the enrichment of omega-3 fatty acids in oils. Lipases are enzymes that catalyze the hydrolysis, esterification, interesterification, acidolysis and alcoholysis reactions (Senanayake and Shahidi, 2000b; Shahidi and Senanayake, 2006; Shimada et al., 2006; Weete et al., 2008). The common
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496 Separation, extraction and concentration processes feature among lipases is that they are activated by an interface. Lipases have been used for many years to modify the structure and composition of foods. Lipases which act on neutral lipids generally hydrolyze the esters of PUFAs at a slower rate than those of more saturated fatty acids (Villeneuve and Foglia, 1997). Use has been made of this relative substrate specificity to increase the concentration of omega-3 PUFAs in seal blubber and menhaden oils by subjecting them to hydrolysis by a number of microbial lipases (Wanasundara and Shahidi, 1998). Concentration of omega-3 fatty acids by enzyme-assisted reactions involves benign reaction conditions and provides an alternative to the traditional concentration methods such as distillation and chromatographic separation (Shahidi and Senanayake, 2006). Furthermore, concentration via enzymatic means may also produce omega-3 fatty acids in the acylglycerol form, which is nutritionally preferred. Studies over the past two decades have used microbial lipases to produce purified omega-3 fatty acids via hydrolysis, esterification or transesterification of marine oils. Tanaka et al. (1992) used six types of microbial lipases (lipases derived from Aspergillus niger, Candida cylindracea, Chromobacterium viscosum, Rhizopus javanicus, Rhizopus delemer and Pseudomonas sp.) to hydrolyze tuna oil and found that the lipase from Candida cylindracea was the most effective one in increasing the DHA content in the non-hydrolyzed fraction of the oil. The DHA content in the non-hydrolyzed fraction was increased three-fold compared with the original oil; however, other lipases were not very effective. Various microbial lipases were evaluated for the enrichment of omega-3 fatty acids from cod liver and sardine oils by selective hydrolysis (Hoshino et al., 1990). The best hydrolysis results were obtained for the lipases from Candida cylindracea and Aspergillus niger, but none of the lipases were able to increase the EPA content of the oil considerably. However, over 50% of the total omega-3 fatty acids were produced when these two lipases were used. A Japanese patent (Noguchi and Hibino, 1984) describes a method based on the discrimination of lipases on DHA and EPA for preparation of omega-3 fatty acids. Ethyl esters from sardine and mackerel were hydrolyzed with several lipases derived from Candida cylindracea, Mucor miehei and Aspergillus niger. Selective hydrolysis afforded ethyl ester concentrates of up to 17% DHA and 25% EPA after separation of the hydrolyzed fatty acids. Transesterification of various fish oil TAGs with a stoichiometric amount of ethanol catalyzed by immobilized Rhizomucor miehei lipase under anhydrous solvent-free conditions resulted in a good separation of EPA and DHA (Haraldsson and Kristinsson, 1998). When FFAs from the various fish oils were directly esterified with ethanol under similar conditions, greatly improved results were obtained. When tuna oil comprising 6% EPA and 23% DHA was transesterified with ethanol, 65% conversion into ethyl esters was obtained. The residual glyceride mixture contained 49% DHA and 6% EPA, with 90% DHA recovery into the glyceride mixture and 60% EPA recovery into the ethyl ester product. When the corresponding tuna oil FFAs were directly esterified with ethanol, 68% conversion was obtained.
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The residual FFAs comprised 74% DHA and only 3% EPA. The recovery of both DHA into the residual FFA fraction and EPA into the ethyl ester product was reported to be very high at 83 and 87%, respectively (Haraldsson and Kristinsson, 1998). Schmitt-Rozieres et al. (2000) studied the enrichment of PUFAs from sardine cannery effluents via enzymatic selective esterification. The sardine canning industry produces vast quantities of effluents that need expensive reprocessing. Their oily component contains EPA and DHA up to 10% each. The author’s goal was to develop a process allowing the recovery of these fatty acids. After the removal of solid particles, proteins, and peptides from the crude effluent, the resultant oil was hydrolyzed and EPA and DHA were enriched by selective enzymatic esterification. Using Lipozyme™, DHA was enriched up to 80%, but no enrichment was observed for EPA. By immobilizing Candida rugosa lipase on Amberlite IRC50 cation-exchange resin, a 30% EPA enrichment was obtained. Shimada et al. (2001) also attempted to purify PUFAs by taking advantage of the enzyme-catalyzed reactions. When FFAs originating from PUFA-containing oil were selectively esterified with lauryl alcohol (LauOH) using a lipase acting on the desired PUFAs very weakly, the PUFA was efficiently enriched in the FFA fraction. In addition, when selective alcoholysis of ethyl esters originating from PUFA-containing oil with LauOH was carried out, the PUFA ethyl ester (EtPUFAs) was enriched to a desired purity in the unreacted ethyl ester fraction. These reaction mixtures contain LauOH, PUFA (EtPUFAs), and lauryl esters, and their molecular weights are different from one another. Hence, PUFA or EtPUFAs can be easily separated by conventional distillation. Selective esterification increased the purity of DHA, GLA, and arachidonic acid (ARA; 20:4n-6) to 91, 98, and 96 wt%, respectively. Selective alcoholysis was also effective for increasing the purity of ethyl docosahexaenoate to 90 wt%. Kojima et al. (2006) studied the enzymatic acidolysis and acylglycerol synthesis using PUFAs with lipases from Pseudomonas fluorescens HU380 (HU-lipase), P. fluorescens AK102 (AK-lipase), and Candida rugosa (CR-lipase). The acidolysis of triolein with EPA or DHA in n-hexane was evaluated with lipases immobilized on Celite 545. HU-lipase showed the highest incorporation rate at a low temperature (10 °C) with either EPA or DHA as the acyl donor, and the rate decreased with increasing reaction temperature. At 45 °C, the rates for EPA and DHA were 7.1 and 0.5 relative to those at 10 °C, respectively. The EPA incorporation rate was even higher at a low temperature (10 °C), and the DHA incorporation rate increased with decreasing temperature. Although AK-lipase showed the reverse tendency for incorporation rate, the DHA incorporation rate increased with increasing reaction temperature with both PUFAs. HU-lipase reacted well with PUFAs such as DHA, EPA, ARA, mead acid (MA), and dihomo-g-linolenic acid (DGLA) on acidolysis and glyceride synthesis. The reactivities of AK-lipase toward these PUFAs except for DGLA, i.e. MA, ARA, EPA, and DHA, were low for both reactions.
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498 Separation, extraction and concentration processes Halldorsson et al. (2003) investigated the use of lipases as catalysts for separating EPA and DHA in fish oil by kinetic resolution based on their fatty acid selectivity. Esterification of FFA from various types of fish oils with glycerol by immobilized Rhizomucor miehei lipase under water-deficient, solvent-free conditions resulted in a highly efficient separation of EPA and DHA. Reactions were conducted at 40 °C with a 10% dosage of the lipase preparation under vacuum to remove the coproduced water, thus rapidly shifting the reaction toward the products. The bulk of the fatty acids, together with EPA, were converted into acylglycerols, whereas DHA remained in the residual FFA. When FFA from tuna oil comprising 5% EPA and 25% DHA were esterified with glycerol, 90% conversion into acylglycerols was obtained. The residual FFA contained 78% DHA and only 3% EPA, with 79% DHA recovery. EPA recovery in the acylglycerol fraction was 91%. The type of fish oil and extent of conversion were highly important parameters in controlling the degree of concentration of EPA and DHA.
17.8 Integrated methods for the concentration and purification of omega-3 fatty acids Most of the separation methods described above, when used alone, can only concentrate and purify omega-3 fatty acids to a limited extent. Therefore, two or more procedures are often required to produce highly purified omega-3 fatty acids. A simple and inexpensive method involving saponification of wet biomass, followed by transmethylation, winterization and urea adduction in a sequential manner has been recently developed for concentration of DHA from Crypthecodinium cohnii biomass (Mendes et al., 2007) (Fig. 17.2). The algal biomass grown in shake flasks is concentrated by centrifugation and the wet concentrate is kept frozen until needed. Fatty acids are extracted by direct saponification of wet biomass with KOH-ethanol. Before extracting unsaponifiable matter with hexane, water is added to shift the equilibrium distribution of unsaponifiable matter to the hexane phase. The hydroalcoholic phase, containing the soaps, is acidified with HCl and free fatty acids are extracted with hexane. The organic phase, containing FFAs, is dried with anhydrous sodium sulfate and the solvent is evaporated to recover FFAs. The FFAs are then methylated using transmethylation reagent and sulfuric acid. The mixture containing methyl esters are then winterized at –18 °C. Subsequently, the liquid fraction is separated from the winterized crystals. The liquid fraction is added to a urea-methanol saturated solution and the urea-fatty acids adducts are crystallized; the crystals are separated by centrifugation and the filtrate (non-urea complexing fraction), which contains DHA, is finally recovered. The highest DHA recovery (49.9%) was obtained at 24 °C at a urea/fatty acid ratio of 4.0, corresponding to 89.4% of DHA of the total fatty acids (Mendes et al., 2007).
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Methods of concentration and purification of omega-3 fatty acids 499 Microalgal wet biomass Saponification Free fatty acids Transmethylation Fatty acid methyl esters (FAMEs) Winterization DHA methyl esters Urea adduction Highly purified DHA methyl esters
Fig. 17.2 Scheme for the DHA concentration and purification of microalgal biomass from Crypthecodinium cohnii.
Medina et al. (1995) described a two-step process, urea adduction followed by preparative high-performance liquid chromatography, for concentration of omega-3 fatty acids from the marine microalga, Isochrysis galbana. By the urea adduction method, a mixture that contained 94% (w/w) stearidonic (SA), EPA, and DHA acids (4:1 urea/fatty acid ratio and 4 °C crystallization final temperature) was obtained from cod liver oil fatty acids. Further purification of SA, EPA, and DHA was achieved with reverse-phase C18 columns. These isolations were scaled up to a semi-preparative column. A PUFA concentrate was isolated from I. galbana with methanol/water (80:20, w/w) or ethanol/ water (70:30, w/w). With methanol/water, a 96% EPA fraction with 100% yield was obtained, as well as a 94% pure DHA fraction with a 94% yield. With ethanol/water as the mobile phase, EPA and DHA fractions obtained were 92% pure with yields of 84 and 88%, respectively. Guil-Guerrero and Belarbi (2001) also used an integrated method to purify DHA and EPA from cod liver oil. The process consisted of four main steps: (i) saponification of the oil, (ii) use of urea inclusion adducts method to obtain PUFAs, (iii) PUFA methylation, and (iv) silica gel column chromatography of the methylated PUFAs. Silica gel chromatography yielded highly pure DHA in the process (100% purity, 64% yield). For EPA, the recovery in the combined process was 29.6%, and the final purity was 90.6%, owing to the simultaneous elution of other PUFA esters. The recovery in the urea adduction method was strongly enhanced by application of orbital agitation during the crystallization process, in which EPA yield increased from 60–70% without agitation to 90–97% at 800 rpm; stearidonic acid (18:4n−3) yield ranged from 60–75% without agitation to 87–95% at 800 rpm, and DHA yield varied from 53–73% without agitation to 85–99% at 800 rpm. Chakraborty and Paulraj (2008) used enzymatic hydrolysis followed © Woodhead Publishing Limited, 2010
500 Separation, extraction and concentration processes by urea adduction to purify EPA and ARA from sardine oil. The enzyme used for the hydrolysis of sardine oil was the lipase derived from Bacillus licheniformis MTCC 6824. The enzyme exhibited more hydrolytic resistance toward the ester bonds of EPA and ARA than those of other fatty acids and was proved to be effective for increasing the concentration of EPA and ARA from sardine oil. Utilizing this fatty acid specificity, EPA and ARA from sardine oil were enriched by lipase-mediated hydrolysis followed by urea fractionation at 4 °C. The purified lipase produced the highest degree of hydrolysis for SFAs and MUFAs (81.5 and 72.3%, respectively, from their initial content in sardine oil) after 9 h. The profile of conversion by lipase catalysis showed a steady increase up to 6 h and thereafter plateaued. Lipasecatalyzed hydrolysis of sardine oil followed by urea adduction with methanol provided FFAs containing 55.4% EPA and 5.8% ARA, respectively, after complexation of saturated and less unsaturated fatty acids. The combination of enzymatic hydrolysis and urea adduction proved to be a promising method to obtain highly concentrated EPA and ARA from sardine oil. Yuzo et al. (2006) employed enzymatic hydrolysis, using lipase from Pseudomonas fluorescens strain HU380, and urea adduction to concentrate EPA and DHA from refined cod oil. The starting oil had 12.2% EPA and 6.9% DHA. Lipase-catalyzed hydrolysis followed by urea adduction provided FFAs with 43.1% EPA and 7% DHA. The resulting yield of concentrated total fatty acids comprised 2.6% of the fatty acids from the cod oil. Thus, EPA was particularly concentrated in the fatty acids derived from refined cod oil as a result of enzymatic hydrolysis followed by urea adduction. On the other hand, hydrolysis of cuttlefish oil with AK-lipase (lipase from Pseudomonas fluorescens strain AK102), followed by urea adduction increased the EPA content from 14.2 to 16.8%, and DHA content from 16.3 to 44.6%. The yield of purified total fatty acids by urea concentrate was 9.4% of the fatty acids from the cuttlefish oil. Thus, DHA was particularly concentrated in this study. A US patent (Zaks and Gross, 1999) disclosed an enzymatic process for preparing an oil-based product significantly enriched in omega-3 fatty acids. The process is a two-step procedure involving lipase-catalyzed transesterification of TAGs followed by low-temperature crystallization. The process yields a mixture of highly pure monoglycerides, at least 60% of which contain omega-3 fatty acids. The process can also be used to produce omega-3 enriched TAG products. Enzymatic methods of enrichment of omega-3 fatty acids have also been investigated under supercritical conditions. Lin et al. (2006) examined the enrichment of omega-3 fatty acids in TAGs of menhaden oil by lipasecatalyzed trans-esterification under supercritical CO2. Before the reaction, menhaden oil was treated by urea adduction to produce 80.1% of omega-3 PUFAs containing 29.4% EPA and 41.8% DHA. Using the sn-1,3-specific lipase from Mucor miehei, the effect of various operating parameters on the reaction was studied, including co-solvent concentration, reaction temperature,
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Methods of concentration and purification of omega-3 fatty acids 501
time, pressure, and substrate ratio (free omega-3 fatty acids: TAGs). Water and ethanol were examined as co-solvents, and both fluids exhibited a maximum for omega-3 fatty acid content in TAGs as a function of concentration. Pressure up to 103.4 bar had a significant positive effect on the conversion of omega-3 PUFA onto TAGs. By further increasing pressure, the conversion rate decreased owing to the transformation of the spatial structure of lipase that leads to deactivation. The enzyme exhibited a good performance and stability in the region of 323 K. The optimal substrate molar ratio of TAG to omega-3 PUFA is about 1:4 taking into consideration the TAG inhibition. The conversion in supercritical CO2 appeared 40% higher than in n-hexane at ambient pressure after 5h. In another study, Lin and Chen (2008) used esterification of free omega-3 fatty acids with enriched omega-6 TAGs to produce a desired structured lipid with an omega-3/omega-6 ratio of 4 by using lipases under supercritical CO2. Comparing four different types of lipases, sn-1,3-specific lipase from Mucor miehei had the highest degree of incorporation under 10 wt% loading amount of total substrates. The optimal operating parameters under 10.2 MPa and 323.15 K supercritical CO2 could attain the desired omega-3/ omega-6 ratio in 6 h. Because of the negative effect on the enzyme activity by the enriched omega-6 TAGs, the optimal substrate ratio of the enriched omega-6 TAGs and the omega-3 fatty acids was chosen as 1/4. To enhance the solubility of omega-3 fatty acids in supercritical CO2, ethanol was applied as a co-solvent and reached an optimal input at 10% of total substrates. The enzyme maintained 81% of its initial activity because of moisture removal from the surface of the enzyme after seven cyclic pressurization/ depressurizations.
17.9 Conclusions Production of purified omega-3 fatty acids from natural source materials may be achieved via a number of techniques, namely urea adduction, chromatography, low-temperature crystallization, supercritical fluid extraction, enzymatic splitting, molecular distillation, as well as a combination of any of the above methods. Purified omega-3 fatty acid products thus produced may be in the form of free fatty acids, alkyl esters or acylglycerols. Each method had its own benefits and drawbacks. Of the methods described above, molecular distillation is highly energy consuming and results in a significant distraction of labile highly unsaturated fatty acids. Other methods, which require the use of organic solvents, include low-temperature crystallization, chromatography and supercritical CO2 extraction, among others. These processes have a number of drawbacks. For example, low-temperature crystallization of acylglycerols typically results in only a small omega-3 enrichment of the product, and chromatography and supercritical fluid extraction are expensive and difficult to scale up. Owing to the potential benefits of having the concentrates in the © Woodhead Publishing Limited, 2010
502 Separation, extraction and concentration processes acylglycerol form, enzymatic procedures have become popular. However, there are several limitations to the enzymatic processes as well. These include the necessity for a complex separation of the product from free saturated fatty acids, use of complex multi-enzymatic systems and the low efficiency that results in an insufficient degree of enrichment. Urea adduction has been very successful in enriching omega-3 fatty acids on an industrial scale and the products thus formed are in the free fatty acid or simple ester forms. Efficient and cost-effective methods of enriching the level of omega-3 fatty acids will continue to be needed in order to reduce the cost and to meet the future demand for highly purified omega-3 products.
17.10 References Ackman R G, Ke P J and Jangaard P M (1973), ‘Fractional vacuum distillation of herring oil methyl esters’, J Am Oil Chem Soc, 50, 1–8. Alkio M, Gonzalez C, Jäntti M and Aaltonen O (2000), ‘Purification of polyunsaturated fatty acid esters from tuna oil with supercritical fluid chromatography’, J Am Oil Chem Soc, 77, 315–321. Bousquet O and Goffic F L (1995), ‘Countercurrent chromatographic separation of polyunsaturated fatty acids’, J Chromatogr, 704, 211–216. Catchpole O J, Grey J B and Noermark K A (2000), ‘Fractionation of fish oils using supercritical CO2 and CO2+ethanol mixtures’, J Supercrit Fluids, 19, 25–37. Chakraborty K and Paulraj R (2008), ‘Enrichment of eicosapentaenoic acid from sardine oil with D5-olefinic bond specific lipase from Bacillus licheniformis MTCC 6824’, J Agric Food Chem, 56, 1428–1433. Chen T and Ju Y (2001), ‘Polyunsaturated fatty acid concentrates from borage and linseed oil fatty acids’, J Am Oil Chem Soc, 78, 485–488. Davarnejad R, Kassim K M, Zainal A and Sata S A (2008), ‘Extraction of fish oil by fractionation through supercritical carbon dioxide’, J Chem Eng Data, 53, 2128– 2132. Du Q, Shu A and Ito Y (1996), ‘Purification of fish oil ethyl esters by high-speed countercurrent chromatography using non-aqueous solvent system’, J Liq Chromatogr Relat Technol, 19, 1451–1457. Gandhi N (1997), ‘Applications of lipase’, J Am Oil Chem Soc, 74, 621–634. Guil-Guerrero J L and Belarbi E (2001), ‘Purification process for cod liver oil polyunsaturated fatty acids’, J Am Oil Chem Soc, 78, 477–484. Haagsma N, Gent C M, Luten J B, Jong R W and Doorn E (1982), ‘Preparation of an w3 fatty acid concentrate from cod liver oil’, J Am Oil Chem Soc, 59, 117–118. Halldorsson A, Kristinsson B, Glynn C and Haraldsson G G (2003), ‘Separation of EPA and DHA in fish oil by lipase-catalyzed esterification with glycerol’, J Am Oil Chem Soc, 80, 915–921. Han D S, Ahn H B and Shin H K (1987), ‘Separation of EPA and DHA from fish oil by solubility differences of fatty acid salts in ethanol’, Korean J Food Sci Technol, 19, 430–434. Haraldsson G G and Kristinsson B (1998), ‘Separation of eicosapentaenoic acid and docosahexaenoic acid in fish oil by kinetic resolution using lipase’, J Am Oil Chem Soc, 75, 1551–1556. Hayashi K and Kishimura H (1993), ‘Preparation of n-3 PUFA ethyl ester concentrates from fish oil by column chromatography on silicic acid’, Nippon Suisan Gakkaishi, 59, 1429.
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Methods of concentration and purification of omega-3 fatty acids 503 Hayes D G (2006), ‘Effect of temperature programming on the performance of urea inclusion compound-based free fatty acid fractionation’, J Am Oil Chem Soc, 83, 253–259. Hayes D G, Alstine J M V and Setterwall F (2000), ‘Urea-based fractionation of seed oil samples containing fatty acids and acylglycerols of polyunsaturated and hydroxy fatty acids’, J Am Oil Chem Soc, 77, 207–213. Hoshino T, Yamane T and Shimuzu S (1990), ‘Selective hydrolysis of fish oil by lipase to concentrate w3-polyunsaturated fatty acids’, Agric Biol Chem, 54, 1459–1467. Hwang L S and Liang J (2001), ‘Fractionation of urea-pretreated squid visceral oil ethyl esters’, J Am Oil Chem Soc, 78, 473–476. Jachmanián I, Margenat L, Torres A I and Grompone M A (2007), ‘Selectivity of supercritical CO2 in the fractionation of hake liver oil ethyl esters’, J Am Oil Chem Soc, 84, 597–601. Kojima Y, Sakuradani E and Shimizu S (2006), ‘Acidolysis and glyceride synthesis reactions using fatty acids with two Pseudomonas lipases having different substrate specificities’, J Biosci Bioeng, 102, 179–183. Kris-Etherton P, Eckel R H, Howard B V, St Jeor S and Bazzarre T L (2001), ‘AHA Science Advisory: Lyon diet heart study. Benefits of a Mediterranean-style, national cholesterol education program/American Heart Association Step I dietary pattern on cardiovascular disease’, Circulation, 103, 1823–1825. Lands, W E M (2003), ‘Diets could prevent many diseases’. Lipids, 38, 317–321. Létisse M and Comeau L (2008), ‘Enrichment of eicosapentaenoic acid and docosahexaenoic acid from sardine by-products by supercritical fluid fractionation’, J Sep Sci, 31, 1374–1380. Létisse M, Rozieres M, Hiol A, Sergent M and Comeau L (2006), ‘Enrichment of EPA and DHA from sardine by supercritical fluid extraction without organic modifier I. Optimization of extraction conditions’, J Supercrit Fluid 38, 27–36. Liang J and Hwang L S (2000), ‘Fractionation of squid visceral oil ethyl esters by shortpath distillation’, A Am Oil Chem Soc, 77, 773–777. Lin, T and Chen, S (2008), ‘Enrichment of n-3 polyunsaturated fatty acids into acylglycerols of borage oil via lipase-catalyzed reactions under supercritical conditions’, Chem Engi J, 141, 318–326. Lin T, Chen S and Chang A (2006) ‘Enrichment of n-3 PUFA contents on triglycerides of fish oil by lipase-catalyzed transesterification under supercritical conditions’, Biochem Eng J, 29 27–34. Mansour M P (2005), ‘Reversed-phase high-performance liquid chromatography purification of methyl esters of C(16)–C(28) polyunsaturated fatty acids in microalgae, including octacosaoctaenoic acid [28:8(n-3)]’, J Chromatogr A, 1097, 54–58. Medina R, Giménez A G, Camacho F G, Pérez J A S, Grima E M and Gómez A C (1995), ‘Concentration and purification of stearidonic, eicosapentaenoic, and docosahexaenoic acids from cod liver oil and the marine microalga Isochrysis galbana’, J Am Oil Chem Soc, 72, 575–583. Mendes A, da Silva T L and Reis A (2007), ‘DHA concentration and purification from the marine heterotrophic microalga Crysthecodinium cohnii CCMP 316 by winterization and urea complexation’, Food Technol Biotechnol, 45, 38–44. Merck Index (1983), An encyclopedia of chemical, drugs and biologicals. 10th edn., Rahway, NJ, Merck and Co., Inc. Mishira V K, Temelli F and Ooraikul B (1993), ‘Extraction and purification of omega-3 fatty acids with an emphasis on supercritical fluid extraction’, Food Res Int, 26, 217–226. Murayama W, Kosuge Y, Nakaya N, Nunogaki Y, Nunogaki K, Cazes J and Nunogaki H (1988), ‘Preparative separation of unsaturated fatty acid esters by centrifugal partition chromatography’, J Liq Chromatogr, 19, 283–300. Noguchi Y and Hibino H (1984), ‘Highly unsaturated fatty acid lower ester concentration and separation process’, Japanese Patent, 59–14793. © Woodhead Publishing Limited, 2010
504 Separation, extraction and concentration processes Perretti G, Motori A, Bravi E, Favati F, Montanari L and Fantozzi P. (2007), ‘Supercritical carbon dioxide fractionation of fish oil fatty acid ethyl esters’, J Supercrit Fluid, 40, 349–353. Schmitt-Rozieres M, Deyris V and Comeau L (2000), ‘Enrichment of polyunsaturated fatty acids from sardine cannery effluents by enzymatic selective esterification’, J Am Oil Chem Soc, 77, 329–332. Senanayake S P J N (2000), ‘Enzyme-assisted synthesis of structured lipids containing long-chain omega-3 and omega-6 polyunsaturated fatty acids’, PhD thesis, St. John’s, NF, Canada, Memorial University of Newfoundland. Senanayake S P J N and Shahidi F (2000a), ‘Concentration of docosahexaenoic acid (DHA) from algal oil via urea complexation’, J Food Lipids, 7, 51–61. Senanayake S P J N and Shahidi F (2000b), ‘Structured lipids containing long-chain omega-3 polyunsaturated fatty acids’, in Shahidi F, Seafood in health and nutrition. Transformation in fisheries and aquaculture: global perspectives, St. John’s, NF, Canada, ScienceTech Publishing Co, 29–44. Shahidi F and Senanayake, S P J N (2006). Nutraceuticals and specialty lipids, in Shahidi F, Nutraceuticals and specialty lipids and their co-products, Boca Raton, FL, CRC Press, 1–25. Shahidi F and Wanasundara U N (1998), ‘Omega-3 fatty acid concentrates: nutritional aspects and production technologies’, Trends Food Sci Technol, 9, 230–240. Shimada Y, Nagao T and Watanabe Y (2006), ‘Application of multistep reactions with lipases to the oil and fat industry’, in Shahidi F, Nutraceuticals and specialty lipids and their co-products, Boca Raton, FL, CRC Press, 365–386. Shimada Y, Sugihara A and Tominaga Y (2001), ‘Enzymatic purification of polyunsaturated fatty acids’, J Biosci Bioeng, 91, 529–538. Stout V F, Niisson W B, Krzynowek J and Schlenk H (1990), ‘Fractionation of fish oil and their fatty acids’, In Stansby M E, Fish oils in nutrition, New York, Van Nostrand Reinhold, 73–119. Stout V F and Spinelli J (1987), ‘Polyunsaturated fatty acids from fish oils’, US Patent 4,675,132. Tanaka Y, Hirano J and Funada T (1992), ‘Concentration of docosahexaenoic acid in glyceride by hydrolysis of fish oils with Candida cylindracea lipase’, J Am Oil Chem Soc, 69, 1210–1214. Teshima S, Kanazawa A and Tokiwa S (1978), ‘Separation of polyunsaturated fatty acids by column chromatography on silver nitrate-impregnated silica gel’, Bull J Soc Sci Fish, 44, 927. Temelli F, Leblance E and Long F (1995), ‘Supercritical CO2 extraction of oil from Atlantic mackerel (Scomber scombrus) and protein functionality’, J Food Sci, 60, 703–706. United States Food and Drug Administration (2004). ‘FDA announces qualified health claims for omega-3 fatty acids’. Press release. http://www.fda.gov/bbs/topics/news/2004/ NEW01115.html. Retrieved 09-10-2009. Villeneuve P and Foglia T A (1997), ‘Lipase specificities: Potential application in lipid bioconversions’, Inform, 8, 640–650. Wanasundara U N (1996), ‘Marine oils: stabilization, structural characterization and omega-3 fatty acid concentration’, PhD thesis, St. John’s, NF, Canada, Memorial University of Newfoundland. Wanasundara U N and Fedec P (2002), ‘Centrifugal partition chromatography (CPC): emerging separation and purification technique for lipids and related compounds’, Inform, 13, 726–730. Wanasundara U N and Shahidi F (1998), ‘Lipase-assisted concentration of n-3 polyunsaturated fatty acids in acylglycerols from marine oils’, J Am Oil Chem Soc, 75, 945–951. Wanasundara U N and Shahidi F (1999), ‘Concentration of omega-3 polyunsaturated fatty
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Methods of concentration and purification of omega-3 fatty acids 505 acids of seal blubber oil by urea complexation: Optimization of reaction conditions’, Food Chem, 65, 41–49. Weete J D, Lai O and Akoh C C (2008), ‘Microbial lipases’ in Akoh C C and Min D B, Food lipids: chemistry, nutrition and biotechnology, Boca Raton, FL, CRC Press, 767–806. Yamagouchi K, Murakami W, Nakano H, Konosu S, Kokura T, Yamamoto H, Kosaka M and Hata K (1986), ‘Supercritical carbon dioxide extraction of oils from Antarctic krill’, J Agric Food Chem, 34, 904–907. Yokoyama M, Origasa H, Matsuzaki M, Matsuzawa Y, Saito Y, Ishikawa Y, Oikawa S, Sasaki J, Hishida H, Itakura H, Kita T, Kitabatake A, Nakaya N, Sakata T, Shimada K and Shirato K (2007), ‘Effects of eicosapentaenoic acid on major coronary events in hypercholesterolaemic patients (JELIS): a randomised open-label, blinded endpoint analysis’, Lancet, 369, 1090–1098. Yuzo K, Eiji S and Sakayu S (2006), ‘Different specificity of two types of Pseudomonas lipases for C20 fatty acids with a D5 unsaturated double bond and their application for selective concentration of fatty acids’, J Biosci Bioeng, 101, 496–500. Zaks A and Gross A T (1999), ‘Enzymatic production of monoglycerides containing omega-3 unsaturated fatty acids’, US Patent 5935828. Zuta C P, Simpson B K, Chan H M and Phillips L (2003), ‘Concentrating PUFA from mackerel processing waste’, J Am Oil Chem Soc, 80, 933–936.
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506 Separation, extraction and concentration processes
18 Extraction of natural antioxidants from plant foods E. Conde, A. Moure, H. Domínguez and J. C. Parajó, University of Vigo, Spain
Abstract: An overview of the most studied vegetal sources of compounds with antioxidant properties and their biological action is firstly presented. The conventional and alternative extraction processes are discussed, special emphasis is given to those using biorenewable, environmentally friendly solvents. Selected examples of integrated processes for extraction, concentration and purification of extracts are included. Key words: antioxidants, plants, extraction, purification.
18.1 Introduction The utilization of phytochemicals in the food and cosmetic industries has attracted public and scientific interest, because of their perceived efficiency, low cost and lack of toxicity. The number of recent publications and patents on antioxidants has increased considerably. The demand for natural antioxidants, which are presumed to be safer, has risen owing to concerns about the long-term safety and negative consumer perception of some synthetic antioxidants. Many natural compounds show antioxidant activity and may act as flavorings, colorants, preservatives and reinforcers of endogenous antioxidant systems (Astley, 2003; Valenzuela et al., 2003; Halliwell, 2002; Halliwell and Gutteridge, 1999). There is epidemiological and clinical evidence relating antioxidant-rich diets with a decreased risk of degenerative diseases, reduced morbidity and mortality. However, the safety limits of many natural antioxidants (which can consist of a mixture of several active substances) are unknown, and their lack of toxicity should be confirmed (Galati and
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Extraction of natural antioxidants from plant foods 507 O´Brien, 2004; Pokorny, 2007). The biological activity and bioavailability of natural antioxidants are the subject of much research. Edible vegetals and agroindustrial residues are considered to be abundant and promising sources of natural antioxidants. The extraction method, of crucial importance for both technical and economic reasons, ideally should be non-destructive, time efficient and suitable for producing high quantities of extracts, which should be processed by selective techniques to yield concentrates of enhanced antioxidant capacity. This chapter presents an overview of the most studied vegetal sources of antioxidants and extraction processes (in particular those using biorenewable, environmentally friendly solvents). Selected examples of integrated processes for extraction, concentration and purification of extracts are also included.
18.2 Antioxidant activity in food systems According to a classic definition, an antioxidant is any substance that, when present at low concentrations compared with those of an oxidizable substrate (such as lipids, proteins, DNA or carbohydrates), significantly delays or prevents oxidation of that substrate (Halliwell, 2002; Halliwell and Gutteridge, 1999). In food systems, the term antioxidant is used to designate the inhibitors of lipid peroxidation, whereas in biological systems it usually refers to protection of lipids, proteins and DNA against oxidative damage by processes or reactions involving reactive oxygen and nitrogen species (ROS and RNS, respectively). Antioxidant activity and antioxidant capacity, although often used interchangeably, have different meanings. Activity refers to the rate constant of the reaction between the antioxidant and the oxidant species, and capacity refers to the amount (in moles) of a given free radical scavenged by a sample. The basic mechanism of lipid peroxidation (Frankel, 2005; Laguerre et al., 2007; Pokorny, 1991) is presented in Fig. 18.1. Lipid peroxidation is induced by oxygen in the presence of initiators such as heat, free radicals, light, photosensitizing pigments and metal ions. Once these free radicals are formed, lipid peroxidation progresses at a high rate by a radical chain reaction, and oxidation ends by producing secondary non-radical compounds. In this scheme, antioxidants can act at different levels: ∑ ∑ ∑ ∑ ∑
scavenging species responsible for oxidation initiation or preventive antioxidants; interruptors of the propagation of the autoxidation chain reaction or chain-breaking antioxidants; singlet oxygen quenchers; synergists or compounds increasing the activity of chain-breaking antioxidants in a mixture; reducing agents;
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508 Separation, extraction and concentration processes Initiation Formation of radicals: peroxyl (RO2•), alkoxy (RO•) or alkyl (R•) Propagation R• + O2 Æ RO2• RO2• + RH Æ ROOH + R• RO• + RH Æ ROH + R•
ROOH Æ RO• + •OH 2 ROOH Æ RO2• + RO• + H2O Termination 2 R• R• + RO2• Æ Stable products 2 RO2•
Fig. 18.1 Mechanism of lipid oxidation.
∑ ∑
metal chelators, which stabilize metal pro-oxidants (iron or copper cations); inhibitors of specific oxidative enzymes (especially lipoxygenases).
Antioxidants often act by several mixed and co-operative mechanisms. An antioxidant can behave as a pro-oxidant depending on the structure, chemical environment and operational conditions: for example, typical antioxidants such as flavonoids, a-tocopherol or ascorbic acid can act as pro-oxidants in the presence of transition metal ions, whereas carotenoids show pro-oxidant activity at high oxygen pressure (Galati and O’Brien, 2004). Table 18.1 lists representative in vitro tests employed in the assessment of antioxidant properties. The need for developing simple and reliable in vitro analytical antioxidant tests is widely acknowledged, and the convenience of a common standard has been claimed. A mixture of caffeic acid, catechin and epigallocatechin3-gallate, hesperetin, and morin was proposed for this purpose (Luthria and Vinyard, 2008). The problems associated with the determination of antioxidant activity were summarized in several critical reviews (Andrade et al., 2008; Becker et al., 2004; Decker et al., 2005; Fernandez-Panchón et al., 2008; Frankel and Finley, 2008; Gordon, 2001; Halliwell, 2002; Karadag et al., 2009; Kiokias et al., 2008; Magalhães et al., 2008; Moon and Shibamoto, 2009; McDonald-Wicks et al., 2006; Prior et al., 2005; Roginsky and Lissi, 2005; Laguerre et al., 2007; Sánchez-Moreno, 2002; Singh and Singh, 2008; Strube et al., 1997; Verhagen et al., 2003). The antioxidant activity depends on a number of factors, including type of substrate, medium, and initiators; oxidation conditions, partitioning properties of the antioxidant between lipid and aqueous phases, and the experimental method used (Antolovich et al., 2002; Decker et al., 2005; Frankel and Meyer, 2000; Gordon, 2001; Yanishlieva, 2001). Pitfalls of in vitro antioxidant research include the low stability of antioxidants, which can lead to formation and accumulation of oxidation products (Haenen et al., 2006; Knasmüller et al., 2008). © Woodhead Publishing Limited, 2010
Table 18.1 In vitro methods usually employed for testing the antioxidant activity of pure compounds and extracts. The activities considered include antiradical, anti-lipoperoxidation, metal chelation and reducing capacity Antioxidant activity assay
Substrate/initiator
Scavenging, decolorization assay, indirect competition method of chain lipid peroxidation (chain-breaking)
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bC (b-carotene bleaching)
ROO˙
Scavenging, chain-breaking, competition method
CAT (conjugated autoxidizable triene)
ROO˙
Chain-breaking
Crocin bleaching
ROO˙
Scavenging, chain-breaking, competition method
CUPRAC (cupric ion reducing antioxidant activity)
Cu2+
Reducing capacity 2+
Deoxyguanosine oxidation
HO˙, H2O2, Fe
Deoxyribose oxidation
HO˙, H2O2, Fe2+
Scavenging, chain-breaking, chelating Scavenging, chain-breaking, chelating
DMDP
DMDP˙+ (N,N-dimethyl-pphenylenediamine)
Scavenging, decolorization assay, indirect competition method of chain lipid peroxidation (chain-breaking)
DPPH
DPPH˙ (a,a-diphenyl-bpicrylhydrazyl)
Scavenging, decolorization assay, indirect competition method of chain lipid peroxidation (chain-breaking)
FIC (ferrous ion-chelating assay)
Fe2+, Cu2+
Chelating
FRAP (ferric-reducing antioxidant power)
Fe3+-TPTZ
Reducing capacity, indirect competition method of chain lipid peroxidation (chain-breaking)
FTC (antioxidant activity in the linoleic acid system with ROO˙ ferrothiocyanate reagent)
Scavenging, chain-breaking
HOSC (hydroxyl radical scavenging capacity)
Scavenging, chain-breaking
HO˙
HPS (hydrogen peroxide scavenging)
H2O 2
Scavenging, chain-breaking
Hypochlorous acid scavenging
HOCl
Scavenging, chain-breaking
Extraction of natural antioxidants from plant foods 509
ABTS or TEAC (Trolox equivalent antioxidant capacity) ABTS·+ [2, 2¢-azinobis(3ethylbenzothiozoline-6sulfonate)]
Mechanism
Antioxidant activity assay
Substrate/initiator
Mechanism
Liposome oxidation
ROO˙, Fe2+
Scavenging, chain-breaking, chelating
CD (conjugated dienes)
LDL
Oxidation by copper, aldehydes formation
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In vitro LDL oxidizability
LDL
Oxidation by copper, conjugated dienes determination
Iodometric assay of lipid peroxides
LDL
Iodide oxidation by lipid peroxides
TBARS (thiobarbituric acid-reacting substances)
LDL, ROO˙
Adduct malondialdehyde–thiobarbituric acid (MDA– TBA) formation, scavenging, chain-breaking
ORAC (oxygen-radical absorbance capacity)
Oxygen radical, ROO˙
Scavenging, chain-breaking
Peroxyl radical scavenging
ROO˙
Scavenging, chain-breaking
Peroxynitrite scavenging
ONOO–
Scavenging, chain-breaking
2 –
PHOTOCHEM (photochemiluminescence)
O˙
Scavenging, chain-breaking
PM (phosphomolybdenum method)
Mo6+
Reducing capacity
Reduction of the Fremy’s radical
Fremy’s (potassium nitrosodisulfonate)
Scavenging, indirect competition method of chain lipid peroxidation (chain-breaking)
RP (reducing power)
Fe3+
Reducing capacity
SRS (superoxide radical scavenging)
O 2˙ –
Scavenging, chain-breaking
TAR (total antioxidant reactivity)
ROO˙
Scavenging, indirect competition method of chain lipid peroxidation (chain-breaking)
TOSC (total oxidant scavenging capacity)
ROO˙, O2˙–, HO˙, HOCl, LO(O) , ONOO–, 1O2
Scavenging, chain-breaking
TRAA (a-tocopheroxyl radical attenuating ability)
a-Tocopheroxyl
Scavenging
TRAP (total radical-trapping antioxidant parameter)
ROO˙
Scavenging, chain-breaking
Trichloromethyl peroxyl scavenging
CCl3O2˙
Scavenging
510 Separation, extraction and concentration processes
Table 18.1 Continued
Extraction of natural antioxidants from plant foods 511 The need to use a variety of antioxidant activity assays is generally agreed, as the results obtained in individual tests can be contradictory. Some recommendations on the in vitro determination of antioxidant activity include: identification and determination of active compounds; evaluation of protection against oxidation in foods or physiological model systems under conditions resembling the real chemical, physical, and environmental conditions in the systems to be protected; detailed understanding of the oxidation mechanisms; measurement of both initial and secondary products, and utilization of reagent concentrations that are physiologically relevant. Low oxidation levels should also be determined, and possible mechanism modification under accelerated oxidation conditions should be considered (Antolovich et al., 2002; Choe and Ming, 2006; Collins, 2005; Decker et al., 2005; Frankel, 1993; Frankel and Meyer, 2000; Frankel and Finley, 2008; Gordon, 2001; Prior et al. 2005; Huang et al., 2005; Schlesier et al., 2002). Several classifications of antioxidants assays have been proposed, based on mechanisms such as the ability to quench free radicals by hydrogen donation, the ability for electron transfer, or a combination of both (Huang et al., 2005; Prior et al. 2005; MacDonald-Wicks et al., 2006). Results from simple in vitro experiments are difficult to extrapolate to the heterogeneous conditions of multifaceted in vivo systems (Haenen et al., 2006), where the possible action on the target tissues is affected by absorption, distribution, metabolism and excretion. The animal model used and a battery of well-validated tests to assess the broad diversity of oxidative damage and antioxidative defence parameters, are crucial for antioxidant research in vivo. Measurement of total antioxidant capacity was proposed as an integrated parameter to consider the balanced action of all antioxidants and their synergistic interactions in plasma and body fluids (Aruoma, 2003; Ghiselli et al., 2000; Halliwell, 2009; Sies, 2007). However, pitfalls and serious limits of applicability of this concept have been manifested (Frankel and Finley, 2008; Sies, 2007) as well as the strong dependency on the analytical method used (Fernández-Panchón et al., 2008).
18.3 Natural compounds with antioxidant activity and major sources 18.3.1 Natural compounds Natural antioxidants are synthesized by plants, micro-organisms, fungi, and animals. The most important groups of natural antioxidants are listed below. Phenolics refer to monomeric, oligomeric or polymeric compounds with an aromatic ring bearing one or more hydroxyl substituents and functional derivatives, such as esters, methyl ethers and glycosides. Phenolics (including
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512 Separation, extraction and concentration processes simple phenols, coumarins, flavonoids, stilbenes, lignans, hydrolyzable and condensed tannins, and phlorotannins), have powerful antioxidant activities in vitro (von Gadow et al., 1997; Miller and Ruiz-Larrea, 2002; Tabart et al., 2009) and play a role in signals between plants, plant defense against predation (by micro-organisms, insects and herbivores) or in response to environmental stress (such as air pollution, heavy metal ions and UV-B radiation). Simple phenolics include hydroxybenzoic and hydroxycinnamic acids. The main subclasses of flavonoids are anthocyanins, flavanols, flavanones, flavonols, flavones and isoflavones, and anthocyanidins and their glycosides. Tannins are polyphenolic compounds with varying molecular masses. Plant tannins include hydrolyzable tannins (gallotannins, which yield glucose and gallic acid upon hydrolysis, or ellagitannins, which produce ellagic acid), and condensed tannins (proanthocyanidins). Algae contain phlorotannins, formed by the polymerization of phloroglucinol (Singh and Bharate, 2006). Table 18.2 summarizes the phenolic composition of some plants. Terpenes are a large and diverse class of lipophilic secondary plant metabolites made up of isoprene units, and are classified into hemi-, mono(C10), sesqui- (C15), di- (C20), sester- (C25), tri- (C30), and tetraterpenoids (carotenoids), having eight isoprenoid C 5 residues. Monoterpenes, sesquiterpenes and diterpenes are the main components of essential oils, which also contain oxygenated derivatives and other compounds (including aldehydes, ketones, phenolic, acetates and oxides). Antioxidant potential has been reported for terpenes (Escuder et al., 2002; Grassmann, 2005; Grassmann et al., 2002) as well as synergistic effects with phenolics (Milde et al., 2004). Carotenoids (carotenes and xanthophylls) have received attention because of their provitamin and antioxidant roles (Bohm et al., 2002). Vitamin E includes a family of tocopherols and tocotrienols and some of their ester derivatives. Tissues and seeds of plants are the major sources of vitamin E, which is associated with membrane lipids or appears in lipid storage stuctures. Vitamin E is the most important natural antioxidant in vegetable oil-derived foods from rice bran, palm, and wheat germ (Weber and Rimbach, 2002). The richest source is a by-product of soybean processing (‘oil deodorizer distillate’). Vitamin E effectively inhibits the peroxidation of lipids by peroxyl radical scavenging, and shows relevant activities related to the regulation of enzymes and gene expression (Brigelius-Flohé, 2009). Small proteins and peptides isolated from various protein hydrolyzates show antioxidant activity (Chen et al., 2006; Kitts, 2005; Kitts and Weiler, 2003). Their low-molecular-weight (low-MW) and simple structure, easy absorption, stability under different conditions and the lack of immunoreaction have been cited as advantages (Chen et al., 1996). Some peptides from hydrolyzed food proteins exert antioxidant activities against enzymatic and non-enzymatic peroxidation of lipids and essential fatty acids (Chen et al., 2006; Hook et al., 2001). Certain hydrophobic amino acids and their derivatives present high activity, and show synergistic effects with non-peptidic antioxidants
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Extraction of natural antioxidants from plant foods 513 Table 18.2 Phenolic contents (mg g–1 dry weight) of some plants Plant (Latin name)
Part Phenolic compounds
Achyrocline F satureioides Anethum graveolens F Averrhoa carambola L B Capparis spinosa L
CA, 4CQA, 5CQA, 3,4dCQA, 3,5dCQA, 4,5dCQA pAA, Ap, BA, CA, C, EC, CGA, pCouA, PCys, GA, GeA, K, Lt, Myr, Q, SiA CGA, GA, Q CGA, GA CA, pCouA, FeA, Q, R, VA
TPC/TF
145/48** 78.3 30.1 15.9
Cassine orientalis
L
Castanea vulgaris
L
C, EC, EGC, PCyB2, GA, K, 24/24 Myr, Q Ap, FeA, GeA, Ng, Q, R 11.9
Casuarina equisetifolia Cinnamomun zeylanicum Coffea macrocarpa
B L B
CA, FeA, GA, Q K, Q, R CA, FeA, GA, Q
72.1 57.7 48.9
L
24/18
Diospyros neraudii
L
Eugenia tinifolia
L
Ficus microcarpa
B
Geranium purpureum Ilex paraguariensis
L
Monimiastrum acutisepalum Moringa oliefera
L L
C, EC, ECG, EGC, EGCG, PCyB1, PCyB2, GA, K, Q C, EC, ECG, EGCG, PCyB1, K, Q C, EC, EGC,EGCG, PCyB1, PCyB2, GA, Q, K Cat, Cou, pPGu, pVGu, OlA, pPPh, PCA, PR, Syr, Sy, V, iVA HBA, CA, C, EC, pCouA, GeA, Q, R, SyA, VA CA, 4CQA, 5CQA, 3,4dCQA, 3,5dCQA, 4,5dCQA, 3FQA, 4FQA, 5FQA C, EC, ECG, EGC, PCyB1, GA,Q, K CGA, EA, FeA, GA, Q
Myonima obovata
L
Phytolacca americana Pimpinella anisum
L
L
S
C, ECG, EGC, PCyB1, PCyB2, K, Q HBA, CA, pCouA, FeA, R, VA 4CQA, 5CQA, 3,4dCQA, 3,5dCQA, 4,5dCQA, 3FQA, 4FQA, 5FQA
75/19 23a/17a 237/6.3*
Reference Marques and Farah, 2009 Shyu et al., 2009 Prakash et al., 2007b Proestos et al., 2006 Soobrattee et al., 2008 Proestos et al., 2006 Prakash et al., 2007b Prakash et al., 2007b Soobrattee et al., 2008 Soobrattee et al., 2008 Neergheen et al., 2006 Ao et al., 2008
28.2
Proestos et al., 2006 Marques and Farah, 2009
10a/9a
Neergheen et al., 2006 Prakash et al., 2007b Neergheen et al., 2006 Proestos et al., 2006 Marques and Farah, 2009
32.9 31a/32a 9.2
© Woodhead Publishing Limited, 2010
514 Separation, extraction and concentration processes Table 18.2 Continued Plant (Latin name)
Part Phenolic compounds
Ricinus communis
L
EC, EA, GA, GeA, Q, R
Rubus ulmifolius Ruta graveolens
L
Salix aegyptiaca
C
CA, 4CQA, 5CQA, CGA, FeA, KGu, KCouGPy, KCGPy, QGu HBA, CA, C, GeA, FeA, Q, R CA, C, EGCG, pCouA, GA, R, V CA, C, EGCG, pCouA, GA, Myr, V CA, C, EGCG, pCouA, GA, Myr, Q, R, V CA, pCouA, GeA, Lt, Q, VA
L B Spartium junceum
F
Styrax officinalis
L
Syzygium glomeratum Juglans regia
L
Achillea millefolium
TPC/TF
L H L ST
Reference Singh et al., 2009 Dall’Acqua et al., 2008
4.3
Proestos et al., 2006 107/351** Enayat and Banerjee, 2009 64/165** 212/479** 4.8
HBA, CA, C, EC, CouA, GA, 18.4 GeA, FeA, Ng, Q, VA C, EC, ECG, EGCG, EGC, 84a/39a PCyB1, PCyB2, GA, Q, K 5CQA, 3CouQA, 4CouQA, CouA, QA, QGal, QP-d, QR, QX Ap, Apg, CGA, Lt, Ltdg, Ltg, R, Vi Ap, Apg, CGA, Lt, Ltdg, Ltg, R, Vi Ap, Apg, CGA, Lt, Ltdg, Ltg, R, Vi
Proestos et al., 2006 Proestos et al., 2006 Neergheen et al., 2006 Pereira et al., 2007 Benetis et al., 2008
B, bark; C, catkins; F, flowers; H, herbs; L, leaves; S, seeds; ST, stems. a Composition of phenols (mg g–1 fresh weight) of some plants. TF, total flavonoid expressed as mg quercetin equivalent g–1 on a dry weight basis. * Total flavonoid expressed as mg rutin equivalent g–1 on a dry weight basis. ** Total flavonoid expressed as mg catechin equivalent g–1 on a dry weight basis. TPC: total phenolic content expressed as mg gallic acid equivalent g –1 on a dry weight basis. 3,4dCQA, 3,4-dicaffeoylquinic acid; 3,5dCQA, 3,5-dicaffeoylquinic acid; 3CouQA, 3-coumaroylquinic acid; 3FQA, 3-feruloylquinic acid; 4,5dCQA, 4,5-dicaffeoylquinic acid; 4CouQA, 4-coumaroylquinic acid; 4CQA, 4-caffeoylquinic acid; 4FQA, 4-feruloylquinic acid; 5CQA, 5-caffeoylquinic acid; 5FQA, 5-feruloylquinic acid; Ap, apigenin; Apg, apigenin-7-O-glycoside; BA, benzoic acid; C, catechin; CA, caffeic acid; Cat, catechol; CGA, chlorogenic acid; Cou, coumaran; EA, ellagic acid; EC, epicatechin; ECG, epicatechin-3-gallate; EGC, epigallocatechin; EGCG, epigallocatechin-3-gallate; FeA, ferulic acid; GA, gallic acid; GeA, gentisic acid; HBA, p-hydroxybenzoic acid; iVA, isovanillic acid; K, kaempferol; KCouGPy, kaempferol-3-O-(6≤-p-coumaroyl)-b-d-glucopyranoside; KCGPy, kaempferol-3-O-(6≤-caffeoyl)-b-d-glucopyranoside; KGu, kaempferol-3-O-glucuronide; Lt, luteolin; Ltdg, luteolin-3¢,7-di-O-glycoside; Ltg, luteolin-7-O-glycoside; Myr, myricetin; Ng, naringenin; OlA, oleanolic acid; pAA, p-anisic acid; PCA, protocatechuic acid; pCouA, p-coumaric acid; PCyB1, procyanidin B1; PCyB2, procyanidin B2; PCys, proanthocyanidins; pPGu, p-propylguaiacol; pPPh, p-propylphenol; PR, 4-n-propylresorcinol; pVGu, p-vinylguaiacol; Q, quercetin; QA, quercetin 3-arabinoside; QGal, quercetin 3-galactoside; QGu, quercetin-3-O-glucuronide; QP-d, quercetin-3pentoside derivate; QR, quercetin 3-rhamnoside; QX, quercetin 3-xyloside; R, rutin; SiA, sinapic acid; Sy, syringol; SyA, syringic acid; Syr, syringaldehyde; V, vanillin; VA, vanillic acid; Vi, vicenin-2.
© Woodhead Publishing Limited, 2010
Extraction of natural antioxidants from plant foods 515 such as phenolic compounds (Erdmann et al., 2008). Antioxidant peptides have been obtained from soybean (Pyo and Lee, 2007), soy-fermented foods (Gibbs et al., 2004), soy protein fractions (Moure et al., 2006), barley hordein (Chiue et al., 1997), and potato (Wang and Xiong, 2005). Different hydrolysis conditions result in peptide mixtures with different properties (De Mejia and De Lumen, 2006). Food-derived bioactive peptides from animal or plant proteins have regulatory functions in humans after being released in vitro or in vivo in the gastrointestinal tract (De Mejia and De Lumen, 2006). A number of beneficial health effects have been claimed for food-derived products containing bioactive peptides, including blood pressure-lowering effects, cholesterol-lowering ability, antithrombotic and antioxidant activities, enhancement of mineral absorption and/or bioavailability, and cyto- or immunomodulatory effects. Maillard reaction products (MRP) such as Schiff bases, premelanoidins and melanoidins are formed by reactions involving the condensation of the carbonyl group of reducing sugars with the amino group of amino acids and proteins. Polyphenols, ascorbic acid and other carbonyl compounds can participate in the Maillard reaction. MRP are naturally formed during food processing and storage. Their characteristics and MW depend on both the type of reagents and the processing conditions (temperature, time, pH, water activity) (Jing and Kitts, 2004). The antioxidant properties of MRP are positively correlated with color when the mechanisms responsible for the formation of antioxidants and color follow the same pathway (Manzocco et al., 2001). Formation of potentially harmful substances (acrylamide and hydroxymethylfurfural) increases rapidly with temperature and time (Morales et al., 2009). The antioxidant capacity of MRP formed from model compounds has been reported (Borrelli et al., 2003; Chen and Kitts, 2008; Dittrich et al., 2003; Kim and Lee, 2008; Morales and Babbel, 2002). Maillardderived antioxidants increase the plasma tocopherol, decrease the plasma thiobarbituric acid reactive substances (TBARS) concentration (Somoza, 2005), prevent low density lipoprotein (LDL) oxidation (Mesa et al., 2008), show reducing activity (Yilmaz and Akgun, 2008) and radical scavenging capacity (Michalska et al., 2008), and may be employed as multifunctional ingredients (Lindenmeier et al., 2002). Carbohydrates including polysaccharides, such as dextran, pullulan, mannan, and lipopolysaccharide, exhibit higher free radical scavenging activity than their constituent sugars, although lower than Trolox™ (Tsiapali et al., 2001). Although glucans are weak free radical scavengers in solution, they stimulate free radical activity in a murine macrophage cell line and modestly augment the generation of free radicals. Arabinoglucogalactan from Panax noto ginseng roots exhibited a, a-diphenyl-b-picrylhydrazyl (DPPH) free radical scavenging activity (Wu and Wang, 2008), whereas porphyran, a sulfated polysaccharide isolated from Porphyra, delayed the aging process in mice by enhancing the amount of antioxidative enzymes and reducing the risk of lipid peroxidation (Zhang et al., 2003). Fucoidans,
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516 Separation, extraction and concentration processes sulfated algal polysaccharides made up of fucose as the major component, exhibit antioxidant and biological properties (anticoagulant, anti-inflamatory, antiviral, or antitumoral activities) (Zhao et al., 2008). 18.3.2 Sources The number of potential sources of antioxidant compounds is increasing. Table 18.3 summarizes the composition of some examples of the major vegetal sources of bioactive compounds. Food by-products have the additional advantages of being renewable, widely distributed, largely available and inexpensive. Waste streams from the processing of agricultural or industrial feedstocks are particularly attractive as sources of antioxidants, particularly when they are concentrated and create environmental problems (Balasundram et al., 2006; Moure et al., 2001; Obied et al., 2005). Typical wastes employed for antioxidant production and composition of some representative cases are summarized in Tables 18.4 and 18.5. Table 18.3 Phenolic composition of fruits, vegetables and cereals (% oven-dried substrate) Vegetal source Fruits Acerola (Malpighia emarginata DC.) Andean blackberry (Rubus glaucus Berth.) Blueberry (Vaccinium species) Brazilian mango (Mangifera indica L.) Capulí cherry (Prunus serotina Ehrh. var. Capulí). Grape (Vitis vinifera L.)
Phenolic compounds
Reference
CGA, EC, EGCG, PCyB1, R
Mezadri et al., 2008 pCouA-d, CyG, EA-d, EC, Et, GA, Gs, Vasco et al., PCYs, PeG, Q-d 2009 pHBA, C, CA, CGA, pCouA, Cy, De, Taruscio et al., EC, FeA, Myr, Ma, Peo, Pet, Q 2004 KG,Ma, iMa, Q, QAp, QAf, QG, QGal, Ribeiro et al., QR, QX 2008 C, CGA, CyG, EC, K-d, PCYs, Q-d Vasco et al., 2009
C, CA, HCA, oCouA, pCouA, GA, R, Q Kedage et al., 2007 Lemon (Melissa officinalis CA, mCouA, Er7G, He, Het, Na, Ng, Dastmalchi L.) RA et al., 2008 Plum (Prunus salicina CA-d, nCGA, pCouA-d, Cyg, EC, K-d, Vasco et al., Lindl.) (var. ‘Santa Rosa’ PCYs, Q-d 2009 and ‘Beauty’) Portuguese pear (Pyrus Ar, monC, terC, CQA, pCouMA, Ferreira et al., communis L. var. S. monEC, terEC, extEC 2002 Bartolomeu) Strawberry (Fragaria × CyG, EA, K, K-CouG, Pe, PeG, PeR, Zhang et al., ananassa Duch.) HPhAC, Q 2008 CGA, nCGA, pCouQA, CyG, EC, PeR, Usenik et al., Sweet cherry (Prunus avium L.) PeoR, R 2008 © Woodhead Publishing Limited, 2010
Extraction of natural antioxidants from plant foods 517 Table 18.3 Continued Vegetal source
Phenolic compounds
Reference
Vegetables Broccoli (Brassica oleracea L.) Escarole (Cichorium indivia var. latifolium L.) Lettuce (Lactuca sativa var. capitata L.) Onion (Allium cepa)
nCGA, FeGe, PCA, dSiGe, SiFeGe, SiGe, dSiFeGe, SidFeGe dCQA, CGA, CGA-d, CtA, ChiA, KGl, KMG dCQA,CGA, CGA-d, CtA, ChiA, Q-d, QMG FeA, GA, K, PCA, Q
Vallejo et al., 2003 Degl’innoocenti et al., 2008 Degl’innoocenti et al., 2008 Prakash et al., 2007 Reddivari et al., 2007 Degl’innoocenti et al., 2008 Aehle et al., 2004
Potato (Solanum Axa, C, CA, CGA, Cxa, GA, Lu, Vxa tuberosum L.) Rocket salad (Eruca sativa HCAs-d, K-d Mill.) Spinach (Spinacia Pa, Sp oleracea) Cereals Barley (Hordeum vulgare C, CA, pCouA, EC, FeA, GA, PCA, L.) SyA, VA Corn (Zea mays L.) CyG, aCyG, HCAs, pCouA, PCA, He-d, PeG, aPeG, PeoG, aPeoG, Q-d, VA Millets
pHBA, C, CA, CinA, pCouA, FeA, GA,GeA, L, Or, iOr, PCA, SaA, Sap, SiA, Tr, VA, Vi, Vt, iVt Rice (Oryza sativa L.) HBA, CA, CGA, pCouA, FeA, FS, PCA, SiA, SiS, SyA, VA Rye (Secale cereale L.) ALRs, pHBA, CA, FeA, SiA, SyA, VA, VeA Sorghum (Sorghum Apf, Ap, Api, ApiG, MApi, MApiG, bicolor L.) 5MApi, PApi, pHBA, CA, EC, pCouA, CinA, PCyB1, PDe, Er, Er5G, FeA, GA, PCA, GeA, KRG, Lut, Ltn, LtnG, PLtn, Lt, MLtn, MLtnG, 7MLtn, Ng, Ta, TaG, SaA, SiA, SyA, VA Wheat (Triticum aestivum CA, oCouA, pCouA, FeA, SyA, VA L.)
Zhao et al., 2006 Pedreschi and CisnerosZevallos, 2007 Dykes and Rooney, 2006 Tian et al., 2004 Heiniö et al., 2008 Dykes and Rooney, 2006
Mpofu et al., 2006
5MApi, 5-methoxyapigeninidin; 7-MApiG, 7-methoxyapigeninidin-5-glucoside; 7MLtn, 7-methoxyluteolinidin; aCyG, acylated cyanidin-3-glucoside; ALRs, alkylresorcinols; Ap, apigenin; aPeG, acylated pelargonidin-3-glucoside; aPeoG, acylated peonidin 3-glucoside; Apf, apiforol; Api, apigeninidin; ApiG, apigeninidin-5-glu; Ar, arbutin; Axa, antheraxanthin; C, catechin; CA, caffeic acid; CA-d, caffeic acid derivatives; CGA, chlorogenic acid; CGA-d, chlorogenic acid derivates; ChiA, chicoric acid; CinA, cinnamic acid; CQA, caffeoylquinic acid; CtA, caffeoyltartaric acid; Cxa, canthaxanthin; Cy, cyanidin; Cyg, cyanidin glycosides; CyG, cyanidin-3-glucoside; dCQA, dicaffeoylquinic acid; De, delphinidin; dSiFeGe, 1,2¢-disinapoyl-2-feruloylgentiobiose; dSiGe, 1,2-disinapoylgentiobiose; EA, ellagic acid; EA-d, ellagic acid derivatives; EC, epicatechin; EGCG, epigallocatechin gallate; Er, eriodictyol; Er5G, eriodictyol 5-glucoside; Er7G, eriodictyol-
© Woodhead Publishing Limited, 2010
518 Separation, extraction and concentration processes Table 18.3 Continued 7-O-glucoside; Et, ellagitannins; extEC, extension epicatechin; FeA, ferulic acid; FeGe, 1,2diferuloylgentiobiose; FS, 6¢-O-feruloylsucrose; GA, gallic acid; GeA, gentisic acid; Gs, galloyls; HBA, hydroxybenzoic acid; HCA, hydrocaffeic acid; HCAs, hydroxycinnamic acids; HCAs-d, hydroxycinnamic derivates; He, hesperidin; He-d, hesperidin derivates; Het, hesperetin; HPhAC, 3,4,5trihydroxyphenyl-acrylic acid; iMa, isomangiferin; iOr, isoorientin; iVt, isovitexin; K, kaempferol; K-CouG, kaempferol-3-(6¢-coumaroyl)glucoside; K-d, kaempferol derivatives; KG, kaempferol 3-O-glucoside; KGl, kaempferol-3-O-glucuronide; KMG, kaempferol-3-O-(6-O-malonylglucoside); KRG, kaempferol 3-rutinoside-7-glucuronide; L, lucenin-1; Lt, luteolin; Ltn, luteolinidin; LtnG, luteolinidin-5-glucoside; Lu, lutein; Lut, luteoforol; Ma, mangiferin; MApi, 7-methoxyapigeninidin; mCouA, m-coumaric acid; MLtn, 5-methoxyluteolinidin; MLtnG, 5-methoxyluteolinidin-7-glucoside; monC, monomeric catechin; monEC, monomeric epicatechin; Myr, myricetin; Na, naringin; nCGA, neochlorogenic acid; Ng, naringenin; oCouA, o-coumaric acid; Or, orientin; Pa, patuletin; PApi, proapigeninidin; PCA, protocatechuic acid; pCouA, p-coumaric acid; pCouA-d, p-coumaric acid derivatives; pCouMA, p-coumarylmalic acid; pCouQA, p-coumaroylquinic acid; PCyB1, procyanidin B1; PCYs proanthocyanidins; PDe, prodelphinidin; Pe, pelargonidin; PeG, pelargonidin-3-glucoside; Peo, peonidin; PeoG, peonidin 3-glucoside; PeoR, peonidin 3-rutinoside; PeR, pelargonidin-3rutinoside; Pet, petunidin; pHBA, p-hydroxybenzoic acid; PLtn, proluteolinidin; Q, quercetin; QAf, quercetin 3-O-arabinofuranoside; QAp, quercetin 3-O-arabinopyranoside; Q-d, quercetin derivatives; QG, quercetin 3-O-glucoside; QGal, quercetin 3-O-galactoside; QMG, quercetin-3-O(6-O-malonylglucoside); QR, quercetin 3-O-rhamnoside; QX, quercetin 3-O-xyloside; R, rutin; RA, rosmarinic acid; SaA, salicylic acid; Sap, saponarin; SiA, sinapic acid; SidFeGe, 1-sinapoyl-2,2¢diferuloylgentiobiose; SiFeGe, 1-sinapoyl-2-feruloylgentiobiose; SiGe, 1,2,2¢-trisinapoylgentiobiose; SiS, 6¢-O-sinapoylsucrose; Sp, spinacetin; SyA, syringic acid; Ta, taxifolin; TaG, taxifolin-7-glucoside; terC, terminal catechin; terEC, terminal epicatechin; Tr, tricin; VA, vanillic acid; VeA, veratric acid; Vi, violanthin; Vt, vitexin; Vxa, violaxanthin.
Table 18.4 Residual sources proposed for antioxidant manufacture Feedstock
References
Residues from grape and wine
Cruz et al., 2004; Corrales et al., 2008; Lafka et al., 2007; Makris et al., 2007; Pinelo et al., 2006; Arvanitoyannis et al., 2006; Louli et al., 2004
Apple pomace
Schieber et al., 2003
Olive
Obied et al., 2005; McDonald et al., 2001
Husks and hulls
Moure et al., 2000; Conde et al., 2008; Goli et al., 2005; Takeoka and Dao, 2003; Oliveira et al., 2008; DeliormanOrhan et al., 2009; Rubilar et al., 2007
Woods
Pérez-Bonilla et al., 2006; Castro et al., 2008; Moure et al., 2005
Leaves
Ferreira et al., 2007; Ozsoy et al., 2008
Seeds
Liu and Yao, 2007; Siddhuraju and Becker, 2007
Peels
Berardini et al., 2005; Li et al., 2006
Fruits and vegetables and their processed products and by-products have been extensively investigated. Their activity has been ascribed to the presence of phenolic compounds (Cieślik et al., 2006; Heinonen and Meyer, 2002; Macheix et al., 1990; Peschel et al., 2006; Proteggente et al., 2003; Robards et al., 1999; Sakakibara et al., 2003). Abundant research on fruits © Woodhead Publishing Limited, 2010
Extraction of natural antioxidants from plant foods 519 Table 18.5 Chemical composition of solid agro-industrial by-products (% oven-dried substrate) Solid agro-industrial Phenolic compounds by-product Almond hulls
References
CGA, cCGA, nCGA
Takeoka and Dao, 2003 Almond shells HCinAs, pCouA, VA, SyA Moure et al., 2007 Apple, peach and CA, p-CouA, FeA Gorinstein et al., pear peels 2002 Apple peels CA, C, EC, CGA, FeA, Q, QA, QG, QGal He and Liu, 2008 Schieber et al., 2003 Apple pomace C, EC, CGA, pCouA, pCouQA, PCyB2, FeA, Pht, Phx, Phz, Q, QA, QG, QGal, QR, QRu, QX Banana peel/pulp C, EC, CG Someya et al., 2002 Barley husks dHB, pCouA, GA, SyA, VA, V Conde et al., 2008 Quettier-Deleu Buckwheat hulls/ B2D, B2G, EC, ECG, QGal, QRu, Q et al., 2000 flour Chestnut wood 4HB, Con, EA, FeA, GA, Sin, Sc, SyA, Canas et al., 1999 Syra, V, VA, Um Coconut husks 4HBA, FeA Dey et al., 2003 Corn cobs AV, HMCin, iEu, MEu, Gu, Gua, EGu, Garrote et al., 2007a VGu, VPh, Sy, ASy, Syra, V Citrus peel and seed CA, p-CouA, Er, nEr, FeA, He, nHe, Na, Bocco et al., 1998 Nr, SiA Eucalypt wood 3HB, HMCin, Co, dHEu, MEu, Gu, VGu, Garrote et al., 2007b iEu, V, iVA, HVA Grape skin and seeds C, EC, CA, oCouA, pCouA, CyGl, GA, Lafka et al., 2007 PCA, SyA, Ty, HTy, VA Olive tree leaves ApG, C, CA, Di, DiG, Lt, LtG, Ol, R, Ty, Benavente-García HTy, V, VA, Ve et al., 2000 Olive tree pruning dHB, HMCin, Ol, Sy, Syra, Ty, HTy, V, Conde et al., 2009 HV, VA Potato peel GA, CGA, CA Kanatt et al., 2005 Rice husks AV, 4HB, HMCin, VPh, Gu, Gua, VGu, Garrote et al., 2007 Eu, iEu, MEu, Sy, ASy, Syra, V, HVA Sweetpotato leaves CA, CGA, 3,4dCQA, 3,5dCQA, Islam et al., 2002 4,5dCQA, tCQA Winemaking waste dHB, 3HBA, CinA, pCouA, GA, PCA, Cruz et al., 2004 solids SyA, VA, iVA Wood of olive tree De, iEu, MEu, Gu, Sy, Syra, HSyA, Ty, Castro et al., 2008 branches HTy, Ol, V, HV 3,4dCQA, 3,4-di-O-caffeoylquinic acid; 3,5dCQA, 3,5-di-O-caffeoylquinic acid; 3HB, 3-hydroxybenzaldehyde; 3HBA, 3-hydroxybenzoic acid; 4,5dCQA, 4,5-di-O-caffeoylquinic acid; 4-HB, 4-hydroxybenzaldehyde; 4HBA, 4-hydroxybenzoic acid; ApG, apigenin-7-glucoside; ASy, acetosyringone; AV, acetovanillone; B2D, proanthocyanidin B2; B2G, B2-3´-O-gallate; C, catechin; CA, caffeic acid; cCGA, cryptochlorogenic acid; CG, catechin gallate; CGA, chlorogenic acid; CinA,
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520 Separation, extraction and concentration processes Table 18.5 Continued cinnamic acid; Co, coniferol; Con, coniferaldehyde; CYGl, cyanidin glycosides; De, desaspidinol; dHB, 3,4-dihydroxybenzaldehyde; dHEu, dihydroeugenol; Di, diosmetin; DiG, diosmetin-7-glucoside; EA, ellagic acid, EC, epicatechin; ECG, epicatechin gallate; EGu, 4-ethylguaiacol; ER, eriocitrin; Eu, 4-eugenol; FeA, ferulic acid; GA, gallic acid; Gu, guaiacol; Gua, guaiacylacetone; HCinAs, hydroxycinnamic acids; He, hesperidin; HMCin, 4-hydroxy-2-methoxycinnamaldehyde; HSyA, homosyringic acid; HTy, hydroxytyrosol; HV, homovanillyl alcohol; HVA, homovanillic acid; iEu, isoeugenol; iVA, isovanillic acid; Lt, luteolin; LtG, luteolin-7-glucoside; MEu, methoxyeugenol; Na, naringin; nCGA, neochlorogenic acid; nEr, neoeriocitrin; nHe, neohesperidin; Nr, narirutin; oCouA, o-coumaric acid; OL, oleuropein; PCA, protocatechuic acid; pCouA, p-coumaric acid; pCouQA, p-coumaroylquinic acid; PCY-B2, procyanidin B2; Pht, phloretin; Phx, phloretin xyloglucoside; Phz, phloridzin; Q, quercetin; QA, quercetin-3-arabinoside; QG, quercetin-3-glucoside; QGal, quercetin3-galactoside; QR, quercetin-3-rhamnoside; QRu, quercetin-3-rutinoside; QX, quercetin-3-xyloside; R, rutin; Sc, scopoletin; SiA, sinapic acid; Sin, sinapaldehyde; Sy, syringol; SyA, syringic acid; Syra, syringaldehyde; tCQA, 3,4,5-tri-O-caffeoylquinic acid; Ty, tyrosol; Um, umbelliferone; V, vanillin; VA, vanillic acid; Ve, verbascoside;VGu, 4-vinylguaiacol; VPh, 4-vinylphenol.
is available, including citrus (Dastmalchi et al., 2008; Di Majo et al., 2005; Jayaprakasha and Patil, 2007; Tripoli et al., 2007), berry fruits (Szajdek and Borowska, 2008; Vasco et al., 2009; Taruscio et al., 2004; Zhang et al., 2008) and cherries (Piccolella et al., 2008). Results have also been reported for juices (Gardner et al., 2000), teas (Gupta et al., 2008; von Gadow et al., 1997; Wiseman et al., 1997) coffee (Nardini et al., 2002), wines (Burns et al., 2001; Fogliano et al., 1999; Minussi et al., 2003; Sanchez-Moreno et al., 2003; Villaño et al., 2005), and several beverages (Arts et al., 2000; Richelle et al., 2001; Tabart et al., 2009). The potential of aromatic herbs, spices and medicinal plants was first assessed in the 1950s (Chipault et al., 1952). Their major antioxidant components are vitamins, phenolic acids, flavonoids and flavonoid derivatives, terpenoids, carotenoids, phytoestrogens and minerals. Specific compounds from plants include carnosic acid, carnosol, rosmarinic acid, rosmanol, thymol, carvacrol, gingerol-related compounds, curcumins, capsaicin, capsaicinol, and ascorbic acid (Suhaj, 2006; Yanishlieva et al., 2006). Their effectiveness in different substrates and other biological activities have been reviewed (Srinivasan, 2005; Yanishlieva et al., 2006; Zheng and Wang, 2001). The in vitro antioxidant capacity of cereals is significantly correlated with their polyphenol content, except for maize (Decker et al., 2002). During digestion, the antioxidant capacity of cereals is increased, providing a favorable antioxidative environment for the epithelium in the large intestine (Fardet et al., 2008). However, the in vitro antioxidant capacity of cereals is only an approximate reflection of their in vivo antioxidant effect, owing to differences in solubility/bioavailability within the digestive tract and to the metabolism/conjugation with compounds such as polyphenols. About 6000 seaweed species have been identified and they are grouped into green (chlorophytes), brown (pheophytes) and red (rhodophytes) algae.
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Extraction of natural antioxidants from plant foods 521 Seaweeds are traditionally exploited mainly for soluble polysaccharides and direct utilization as foods. The soluble fraction of red seaweeds is mostly composed of sulfated galactans (agar, carrageenans), whereas the brown seaweeds (alginates, fucans, and laminarans) correspond mostly to reserve b-glucans. The recovery of valuable biomolecules, including vitamins, minerals, lipids, polyphenols, proteins, sterols, pigments, tocopherol derivatives and related isoprenoids, is a topic of growing interest (Herrero et al., 2006; Kumar et al., 2008). The most potent antioxidants of algae are phlorotannins and sulfated polysaccharides (Rupérez et al., 2002; Wang et al., 2008; Xue et al., 2001; Ye et al., 2008). A recent work has reported the irradiation-dependent production of butylhydroxytoluene (BHT) by green algae and cyanobacteria (Babu and Wu, 2008). Mushrooms are also a source of phenolic antioxidants (Cheung et al., 2003).
18.4 Biological activities of natural antioxidants Reactive oxygen species (ROS) play a dual role: a certain physiological level of ROS is crucial for the regulation of cell functions (Wang and Yi, 2008), but ROS and other radicals are also involved in diabetes, cancer, liver disease, aging, arthritis, AIDS, macular degeneration, and autoimmune, inflammatory, cardiovascular and neurodegenerative diseases. Antioxidants prevent ROS concentrations from reaching a damaging level and have been considered as a promising therapy for the prevention and treatment of these diseases (Seifried et al., 2007). The results from many intervention trials with antioxidants failed to demonstrate benefits in humans, probably because the reactive species may not be important, the pathologies were too advanced, the doses used may be wrong or the administered antioxidants do not decrease the oxidative damage (Halliwell, 2009). Antioxidant properties alone are not sufficient to explain their biological properties, some of which are indicated in Table 18.6. Although an antioxidant is a redox agent that in the presence of metal ions could act as a pro-oxidant, in vivo most transition metal ions are protein-conjugated and not available to catalyze free radical reactions (Hadi et al., 2007). The most studied compounds for controlling various disorders (including cardiovascular, neurological and neoplastic diseases) are plant-derived polyphenols. The use of the cardioprotective properties of dietary antioxidants against cardiovascular diseases has been studied and some conflicting findings have been found (Seifried et al., 2007). Natural polyphenolic compounds possess antioxidant, vasorelaxant and antihypertensive properties (Depeint et al., 2002; Erlund, 2004; Ghosh, 2005; Ullah and Khan, 2008). Some effects of polyphenols (such as anti-atherogenesis) can be attributed to their antioxidant properties (German and Walzem, 2000). Selected polyphenols
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Biological activities
Plant or extract
Active compounds
Reference
Antiallergic Antiarthritic
Coronopus didymus Cleome gynandra
Antiasthmatic Antibacterial Anticancer Anticoagulant Antidiabetic Antiglycative Antihyperglycemic
Laurencia undulata Cyperus rotundus Prunus serrulata Cirsium japonicum Psidium guajava L. Helichrysum plicatum Psidium guajava L. Punica granatum
Mantena et al., 2005 Narendhirakannan et al., 2005 Jung et al., 2009 Kilani et al., 2008 Lee et al., 2007 Yin et al., 2008 Hsieh et al., 2007 Aslan et al., 2007 Hsieh et al., 2007 Bagri et al., 2009
Antihyperlipidemic
Coronopus didymus Helichrysum plicatum Punica granatum
Flavonoids, tannins Polyphenolic compounds, tannins, anthroquinones, flavonoids Phlorotannins Tannins, flavonoids, coumarins, sterols Phenolic glucosides, flavonoids, anthocyanins Phenolic acids, flavonoids Phenolic compounds Phenolic acids, flavonoids Phenolic compounds Alkaloids, flavonoids, anthocyanins, hydrolyzable tannins Flavonoids, tannins Phenolic acids, flavonoids Alkaloids, flavonoids, anthocyanins, hydrolyzable tannins Flavonoids and phenolics acids Polyphenols, tannins Bioflavonoids, catechins, procyanidins, phenolic acids Monoterpene glycosides Eugenol Curcuminoids Polyphenols Oligomeric procyanidins, catechins, condensed tannins
Souza et al., 2007 Kim et al., 2007 Hasegawa et al., 2008 Panich et al., 2009 Panich et al., 2009 Ahn et al., 2008 Esquenazi et al., 2002
Anti-inflammatory
Antimelanogenic
Antimicrobial
Geranium robertianum L. Uncaria tomentosa Myracrodruon urundeuva Pycnogenol® Eucalyptus globulus Almina galanga Curcuma aromatic Sasa quelpaertensis Nakai Cocos nucifera L.
Mantena et al., 2005 Aslan et al., 2007 Bagri et al., 2009 Amaral et al., 2009
522 Separation, extraction and concentration processes
Table 18.6 Biological activities reported for antioxidant extracts from some plant foods
Pycnogenol® and Ginkgo biloba Hibiscus tiliaceus L.
Antiparkinson Antipyretic Antitumor Antiulcer Antiviral Apoptotic activities Cytotoxic Hepatoprotective
Mucuna pruriens Coronopus didymus Senna alata Myracrodruon urundeuva Cocos nucifera L. Cyperus rotundus Cyperus rotundus Coronopus didymus Commiphora berryi
Polyphenols, procyanidins Polyphenols, carotenoids, tocopherols, flavonoids, anthocyanins, phytosterols Polyphenols Flavonoids, tannins Flavonoids, tannins, polyphenols Polyphenols, tannins Oligomeric procyanidins, catechins, condensed tannins Tannins, flavonoids, coumarins, sterols Tannins, flavonoids, coumarins, sterols Flavonoids, tannins Phenolic compounds, tannins
Decalepis hamiltonii
Phenolic compounds
Ecklonia cava
Polyphenols
Dhanasekaran et al., 2008 Mantena et al., 2005 Pieme et al., 2008 Souza et al., 2007 Esquenazi et al., 2002 Kilani et al., 2008 Kilani et al., 2008 Mantena et al., 2005 Gowri Shankar et al., 2008 Srivastava and Shivanandappa, 2009 Ahn et al., 2008
Morus alba L.
Flavonol glycosides
Katsube et al., 2006
Castanea sativa
Polyphenols
Almeida et al., 2008
Immunomodulatory activity Inhibitor LDL oxidation Skin protector
Križková et al., 2008 Rosa et al., 2006
Extraction of natural antioxidants from plant foods 523
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Antimutagenic
524 Separation, extraction and concentration processes (extracts from grape and wine, resveratrol, genistein, dihydrodaidzeins) relax endothelial function (Shen et al., 2006). Flavonoids can have favorable effects on atherosclerosis. Lipid-lowering activity has been reported for tea flavonoids (Li et al., 2005), a-tocopherol, and b-carotene, which inhibit the oxidation of LDL and the atherosclerotic process (Odeh and Cornish, 1995). Tocopherols show enhanced immune response and regulation of platelet aggregation (Weber and Rimbach, 2002). Protection against coronary heart disease was reported for resveratrol (Szewczuk and Penning, 2004), anthocyanidins, EGCG, and vitamin E (Violi et al., 2006). Biologically active peptides promote diverse activities (antithrombotic, hypocholesterolemic and antihypertensive actions) relevant to cardiovascular health (Erdmann et al., 2008). Inhibitors of angiotensin-converting enzyme (ACE) play an important role in the regulation of blood pressure and fluid and salt balance in mammals. Immunomodulatory activity has also been reported for peptides (Erdmann et al., 2008). Antiallergenic activity has been reported for flavonoids (Das and Rosazza, 2006; Ghosh, 2005; Hodek et al., 2002; McKay and Blumberg, 2006), which may also preserve T-cell-mediated immunity (Horváthová et al., 2001; Strickland, 2001). Bioactive compounds such as flavonoids (Horváthová et al., 2001) and vitamin E protects against cataracts. Vitamin E decreases oxidative stress and the levels of erythrocyte osmotic fragility in patients on dialysis (Uzum et al., 2006) and had beneficial effects on diabetic patients (Levy and Blum, 2007). Epidemiological studies suggested a positive association between a diet rich in fruit and vegetables and a lower incidence of cancer (Ames et al., 1993; Pan et al., 2008a; Ren et al., 2003; Tanaka and Sugie, 2008). The antioxidant properties may only partly explain the antitumor effects of dietary phenolics, which may exert modulatory actions in cells, interfering in the steps leading to the development of malignant tumors (Kanakis et al., 2007). Supplemental antioxidants may help to prevent cancer only in dietdeficient populations or individuals, and antioxidant anticancer agents may affect patients differently according to their health (Vickers, 2007). Although dietary antioxidants have attracted great interest and are considered safe, they have to pass controlled clinical trials for therapeutic or chemopreventive use, and potential toxic flavonoid–drug interactions should be considered (Galati and O’Brien, 2004). Many natural agents have shown potential in bioassays and animal models, and some of them have been selected for ongoing phase I–III clinical trials (Bonfili et al., 2008). An ideal, effective and acceptable agent should be non-toxic in normal and healthy cells, low cost and efficient against multiple cancers, act according to a known mechanism, deserve public acceptance, and be suitable for oral consumption. Potential cancer preventive agents suppress carcinogenesis by several major mechanisms: inhibiting phase I enzymes or blocking carcinogen formation, inducing phase II (detoxification) enzymes, scavenging DNA reactive agents, suppressing the over-expression of pro-oxidant enzymes, modulating hormone homeostasis,
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Extraction of natural antioxidants from plant foods 525 suppressing hyper-cell proliferation induced by carcinogens, inducing apoptosis, counteracting angiogenesis, and/or inhibiting certain phenotypic expressions of preneoplastic and neoplastic cells (D’Archivio et al., 2008; Shklar, 1998; Ren et al., 2003; Pan et al., 2008a). Various dietary antioxidants are under investigation for their anticancer properties, including curcumin (Tsvetkov et al., 2005; Maheshwari et al., 2006), oleuropein (Bonfili et al., 2008), and tangeretin, nobiletin and resveratrol (Narayanan, 2006). Green tea inhibits tumor incidence and multiplicity in various organs, and recent phase I and II clinical trials have been conducted to explore its anticancer effects (Dou, 2008). Flavonoids in human diet may reduce the risk of various cancers (Kanadaswami et al., 2005), including hormone-dependent breast and prostate cancers (Hodek et al., 2002), intestinal neoplasia (Hoensch and Kirch, 2005) and skin cancer (Singh and Agarwal, 2002). Lycopene shows both in vitro and in vivo anticancer activities, possibly through ROS scavenging, up-regulation of detoxification systems, interference with cell proliferation, induction of gap-junctional communication, inhibition of cell cycle progression and modulation of signal transduction pathways (Bhuvaneswari and Nagini, 2005). A beneficial role of vitamin C on cancer has been claimed, although further clinical trials are needed (Verrax and Calderón, 2008). Vitamin E and vitamin E-based compounds are potent antioxidants regulating peroxidation, some forms display apoptotic activity against cancer cells and restore multidrug resistant tumor cells sensitivity to chemotherapeutic agents (Kline et al., 2007; Sylvester, 2007). Among soy peptides, lunasin suppresses transformation of cells induced by carcinogens and viral oncogenes (De Mejia and De Lumen, 2006), and the Kunitz trypsin inhibitor suppresses ovarian cancer cell invasion by blocking urokinase up-regulation (Xiao et al., 2005). Combinations of antioxidant nutrients act synergistically, and their use is proposed to offer a better quality of life (Shklar, 1998), and to present a co-operative action with chemotherapeutic drugs and radiation therapy (Bonfili et al., 2008; Prasad et al., 2002). The nervous system, rich in fatty acids and iron, is vulnerable to free radical generation (Ullah and Khan, 2008). Reactive species are constantly produced in the brain, for example by leakage of electrons from the mitochondrial electron transport chain to generate superoxide radical (O2–). Increased levels of oxidative damage may occur early in the progression of many neurodegenerative diseases (Philips et al., 1993), as oxidative damage may aggravate the accumulation and precipitation of proteins. Epidemiological studies indicate that dietary antioxidants can limit the incidence of neurodegenerative disorders (Morris, 2002). Polyphenols are active agents in neuroprotection owing to their ability to influence and modulate several cellular processes such as signaling, proliferation, apoptosis, redox balance, and differentiation. Their role in neurological disorders has been studied (Dajas et al., 2002; Horváthová et al., 2001; Singh and Singh, 2008). Promising compounds for neuroprotection include polyphenols such as epigallocatechin gallate (EGCG), curcumin, naringenin, extracts of
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526 Separation, extraction and concentration processes blueberries and Scutellaria (Ullah and Khan, 2008), and extracts from green tea (Bastianetto et al., 2006). Flavonoids and their derivatives are reported to inhibit the growth and development of HIV by suppressing acute infection and inhibiting protease, integrase and reverse transcriptase (Critchfield et al., 1996), whereas scutellarin showed anti-HIV activity (Zhang et al., 2005). The overall role of reactive species in chronic inflammatory diseases is not clear. Reactive species may help explain why such diseases increase cancer risk, but ironically these species can sometimes be anti-inflammatory. Hence the use of antioxidants to treat chronic inflammatory diseases may not be as simple as it originally sounded (Halliwell, 2009). Natural antioxidants with anti-inflamatory activity include phenolics (Lomnitski et al., 2000), resveratrol, flavonoids (Ghosh, 2005; Das and Rosazza, 2006; Hodek et al., 2002; Kim et al., 2004), terpenoids (De las Heras et al., 2003), and glucans (Tsiapali et al., 2001). Therapeutic potential for combating bronchial asthma has been reported for polyphenolic extracts from Laurencia undulata, an edible red alga (Jung et al., 2009). Tyrosinase inhibitors can ameliorate cutaneous hyperpigmentary disorders. Their use is becoming increasingly important in the cosmetic industry owing to their skin whitening action and preventive effects, and there is a growing interest in their medicinal and cosmetic applications (Parvez et al., 2007). Some carotenoids (a- and b-carotene and b-cryptoxanthin) are precursors of vitamin A and protect against chemical oxidative damage and against several kinds of cancer, age-related macular degeneration and UV-induced erythema.
18.5 Extraction of natural antioxidants from plant foods and residues Solid–liquid extraction is a heterogeneous operation involving transfer of solutes from a solid to a fluid. Extractable compounds of vegetal origin are a complex mixture of solutes that can be extracted at different rates depending on their location (outer surface, pores and vacuoles). Extraction involves the following sequential steps: (i) transport of solvent from the bulk solution to the external surface of the particle, (ii) solvent penetration and diffusion in the solid matrix, (iii) solubilization of the components, (iv) transport of the solute(s) through the solid matrix, and (v) transport of the solute(s) from the external surface of the solid to the bulk solution. 18.5.1 Conventional solvent extraction The type of solvent is one of the most influential variables on both extraction yield and type of extracted solutes. Methanol, ethanol, and water are widely
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Extraction of natural antioxidants from plant foods 527 employed for extracting phenols. Methanol exhibits the highest capacity to extract phenolics (Pinelo et al., 2005), but shows toxicity, whereas hot water may extract polyphenols without safety concerns. Ethanol–water mixtures are suited to penetrate the hydrophobic areas of the vegetable matrix and help to precipitate soluble proteins, facilitating further processing. Representative examples of extraction processes involving these solvents are summarized in Table 18.7. Temperature is one of the most influential variables affecting the release of phenols. Increased temperatures favor extraction by enhancing both solubility and diffusion. However, temperature cannot be increased indefinitely, because instability of phenolic compounds and denaturation of membranes may take place at temperatures above 50 °C (Cacace and Mazza, 2003). Partially oxidized polyphenols can exhibit higher antioxidant activity than non-oxidized phenols. Thermal treatment of crude extracts can be performed to modify their composition by oxidation, hydrolysis and isomerization, and can result in the formation of compounds with new antioxidant properties. Therefore, heat-induced antioxidants balance the thermal loss of antioxidants, or the overall antioxidant properties can be increased upon heating (Delgado-Andrade and Morales, 2005). This behavior was reported for catechin, resveratrol, grape extract (Pinelo et al., 2005), and citrus peel (Xu et al., 2007). Steam-processing of broccoli may release more bound phenolic acids from the breakdown of cellular constituents (Roy et al., 2009). Heating vegetal products such as sweet corn increased the total antioxidant activity and the phenolic content (Dewanto et al., 2002). Heat treatment of huyou (Citrus paradisi Changshanhuyou) peel increased the free fraction of phenolic acids and decreased the ester, glycoside, and ester-bound fractions (Xu et al., 2007), whereas the antioxidant capacity of citrus peel [measured by the ABTS and ferric-reducing antioxidant power (FRAP) assays] increased with heating time and temperature owing to the increase of lower MW phenolics (benzoic acids and cinnamic acids). High solvent-to-solid ratios favor phenol yields, but a balance between high costs (derived from solvent usage) and absence of saturation effects has to be found (Pinelo et al., 2005). However, operational variables such as temperature, solvent-to-solid ratio, and type of solvent used only affect solutes weakly linked to the cell wall structure and those contained in vacuoles. Extraction of compounds forming part of the cell walls may require other strategies, such as cell wall breaking or enzymatic treatments. Even though conventional extraction with organic solvents is widely used, it has low efficiency and requires large volumes of solvent, long extraction times and high temperature. This could lead to loss of biological activities and decreased selectivity, as well as to problems derived from organic solvent cost and disposal (Robards 2003). Additionally, organic solvents must be completely removed from the exhausted solids. Recently, there has been an increasing interest in using environmentally friendly extraction technologies for producing extracts of high quality and activity. Some of these techniques
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Raw material
Solvent
Solubles yield or Antioxidant activity TEP* (% dry weight)
Reference
Anethum graveolens (flowers)
W
8
DPPH(–)*1; bC(–)*1
Sharififar et al., 2009
Anethum graveolens (flowers)
E
2.84
TEAC(–)*2; DPPH(–)*2,3; RP(–)*2,3; FIC(–)*4; bC(–)*2
EA
8.28
Hordeum vulgare (seeds) Chamaecyparis obtusa (bark)
E (70%)
4.70
TEAC(–) ; DPPH(–) ; RP(–) ; FIC(–) ; bC(–) RP(–)*1,3; DPPH(–)*1,3; FTC(+)*1(=)*3
E
9.07 (GA)
DPPH(–)*2; TEAC(–)*5; RP(–)*5; bC(–)*5
EA
10.28 (GA)
Cirsium japonicum (leaves) Chukrasia tabularis (leaves)
W
24.13
DPPH(–) ; TEAC(–) ; RP(–) ; bC(–) DPPH(–)*3,6; RP(–)*6; HOSC(–)*6
EA W E (95%)
98.35 (GA) 60.88 (GA) 9.68 (PC)
DPPH(–)*7; RP(–)*7; HOSC(+)*7 DPPH(–)*7; RP(–)*7; HOSC(–)*7 DPPH(+)*1; HOSC(+)*1; RP(+)*1; TA(+)*1; LPPO(+)*1
Kaur et al., 2008
E (96%)
17.74
DPPH(–)*1,8; bC(–)*1,8
Moure et al., 2001
W
17.36
EA E W W
1.09 (GA) 1.94 (GA)
DPPH(–)*1,8; bC(–)*1,8 DPPH(–)*1,(+)*6; FIC(–)*3; RP(–)*1 DPPH(–)*1,(+)*6; FIC(–)*3; RP(–)*1; FTC(+)*1 DPPH(–)*1,6; RP(–)*1 DPPH(–)*2,3; RP(–)*2,3; TEAC(–)*2
Dimocarpus longan Lour. (peel) Gevuina avellana (seeds) Kappaphycus alvarezii
Litchi chinenesis (flowers)
9.92
*2
*2
*2,3
*5
*2,3
*5
*4
Shyu et al., 2009
*2
Liu and Yao, 2007 Marimuthu et al., 2008
*5
Yin et al., 2008
Pan et al., 2008b
Kumar et al., 2008
Liu et al., 2009
528 Separation, extraction and concentration processes
Table 18.7 Examples of extraction of plant food antioxidants using conventional solvent extraction. Symbols (+), (–) and (=) indicate higher/ lower/equal activity compared with the reference standard compound
W
1.60
DPPH(–)*1,8; bC(–)*1,8
W
30.96
FTC(+)*6; bX(+)*7(–)*8; DPPH(–)*7,8; SRS(–)*7; HOSC (–)*7,8; HPS(–)*7,8; RP(–)*7,8; FIC(–)*4
E
34.46
FTC(+)*6; bC(+)*7(–)*8; DPPH(–)*7,8; HOSC(+)*7(–)*8; HPS(–)*7,8; RP(–)*7,8; FIC(–)*4
EA
10.61
E
37 (GA)
FTC(+)*6, bC(+)*7(–)*8, DPPH(–)*7,8; HOSC(+)*7(–) *8; HPS(–)*7,8; RP(–)*7,8; FIC(–)*4 DPPH(–)*2,3,5,7,12(+)*1; SRS(+)*2,12; HOSC(–)*1,3,5 (+)*2,12
E
34 (GA)
DPPH(–)*1,2,3,5,7,12; SRS(–)*2,12; HOSC(–)*1,2,3,5,12
E
27 (GA)
DPPH(–)*1,3,7(+)*2,5,12; SRS(–)*2,12; HOSC(–)
W
–
FTC(+)*6; RP(+)*6; SRS(+)*1,6,8; DPPH(–)*5(+)*8; FIC(+)*1,6,8; HPS(–)*1,6,8
W E (50%)
8, 5.4 (GA) 12, 16.8 (GA)
Citrus sinensis
E EA
10, 10.4 (GA) 9.12, 6.43 (C)
Citrus grandis
EA
3.41, 9.72 (C)
Randia echinocarpa W (pulp) Rosa rubiginosa (seeds) E (96%) Smilax excelsa (leaves)
Syzygium cumini (kernel) Syzygium cumini (pulp) Syzygium cumini (seed coat) Urtica dioica (aerial parts)
9.46
*1,2,3,5,12
Ozsoy et al., 2008
Benherlal and Arumughan, 2007 Benherlal and Arumughan, 2007 Benherlal and Arumughan, 2007 Gülçin et al., 2004
Extraction of natural antioxidants from plant foods 529
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1.97
DPPH(–)*3,6, 8,9,10; TBARS(–)*6,8; SRS(–)*11; HOSC(–)*8 Ningappa et al., 2007 DPPH(+)*2,3,8,9,10; TBARS(+)*6,8; SRS(+)*11; HOSC(+)*8; RP(+)*8; FIC(–)*4 DPPH(–)*3,6, 8,9(+)*10; TBARS(+)*6,8; SRS(–)*11; HOSC(–)*8 DPPH(–)*3; ORAC(–)*12; ABTS(–)*3; RP(–)*2; PM(+)*13 Jayaprakasha et al., 2008 *3 *12 *3 *2 *13 Jayaprakasha et al., DPPH(–) ; ORAC(–) ; ABTS(–) ; RP(–) ; PM(+) 2008 *1 Santos–Cervantes bC(–) et al., 2007 *1,8 *1,8 Moure et al., 2001 DPPH(–) ; bC(–)
Murraya koenigii (leaves)
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*Total extractable polyphenols. E, ethanol; EA, ethyl acetate; W, water. C, catechin; GA, gallic acid; PC, pyrocatechol. *1 , butylhydroxytoluene (BHT); *2, (+)-catechin; *3, ascorbic acid; *4, ethylene diamine tetra-acetic acid (EDTA); *5, quercetin;*6, a-tocopherol; *7, gallic acid; *8, butylhydroxyanisol (BHA); *9, curcumin; *10, b-carotene;*11, superoxide dismutases (SOD); *12, trolox; *13, propyl gallate. bC, b-carotene bleaching method; DPPH, a,a-diphenyl-b-picrylhydrazyl radical scavenging assay; FIC, ferrous ion-chelating assay; FTC, antioxidant activity in the linoleic acid system with ferrothiocyanate reagent; HOSC, hydroxyl radical scavenging assay; HPS, hydrogen peroxide scavenging assay; LPPO, lipid peroxidation in peanut oil; ORAC, oxygen-radical absorbance capacity; PM, phosphomolybdenum method; RP, reducing power; SRS, superoxide radical scavenging assay; TA, total antioxidant capacity assay; TBARS, thiobarbituric acid-reacting substances; TEAC, Trolox equivalent antioxidant capacity assay.
530 Separation, extraction and concentration processes
Table 18.7 Continued
Extraction of natural antioxidants from plant foods 531 (enzyme-assisted aqueous extraction, some novel methods including subcritical water, pressurized liquid extraction, supercritical fluid extraction (SFE) with CO2, and ultrasound- and microwave-assisted processes) are considered in the following subsections. 18.5.2 Enzyme-aided extraction Enzyme processing is used as a pretreatment for conventional and alternative solvent extraction technologies, owing to the ability of enzymes to disrupt cell walls. Enzyme-aided extraction is considered more environmentally friendly than chemical treatments, and the further utilization of both residue and extract is possible (Mandalari et al., 2006). Enzyme treatments have been considered for a variety of materials, including algae, fruits, herbs and cereals. Edible and other abundant and underutilized seaweeds have been studied for extracting water-soluble compounds with antioxidant activity. Most antioxidant compounds are soluble in polar solvents, but the large amounts of highly viscous polysaccharides act as a physical barrier. Enzymes can break down the cell walls and storage materials to release both free compounds and polysaccharides attached to polyphenols (Heo et al., 2005). Representative studies of antioxidant extraction from algae biomass using hydrolytic enzymes are summarized in Table 18.8. Synergistic effects between different enzyme activities (including endopeptidases, carbohydrases and proteases) enhance the release of antioxidants (Siriwardhana et al., 2008; Heo et al., 2005). The variables affecting the extraction yield and activity include: type of raw material and pre-conditioning stages, temperature, type and concentration of solvent, and type and concentration of enzyme(s). This latter effect is shown in Fig. 18.2. Incubation results in increased extraction efficacy (Siriwardhana et al., 2008). Thermal degradation of high-MW polysaccharides enables higher extraction of target antioxidants. Some enzymatic hydrolysates are stable at 100 °C (Athukorala et al., 2006; Heo et al., 2005) and the DPPH radical scavenging activity increases with heating time (Siriwardhana et al., 2008). Heating followed by enzymatic hydrolysis enhances the access of endopeptidases and b-glucanase to laminarin and proteins. Increased antioxidant activity at alkaline pH was observed in Hizikia fusiformis, owing to hydrolysis, ion exchange, desulfation and prevention of the formation of the protein–polyphenol complex (Siriwardhana et al., 2008). A simple procedure to reduce the viscosity of porphyrans of (sulfated polysaccharide of Porphyra) various MW, consisting of treatments with hydroxyl radical, has been reported (Zhang et al., 2003). Algal enzymatic extracts show a variety of other biological activities. The enzymatic extract of Ecklonia cava inhibited proliferation of cancer cells strongly and selectively (Athukorala et al., 2006), and showed in vivo anti-inflammatory, inmunomodulatory and antiproliferative activity (Ahn et al., 2008). Highly sulfated fucose and galactose polysaccharides from
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Table 18.8 Examples of enzyme-aided extraction of plant food antioxidants
Algae Ecklonia cava Hizikia fusiformis © Woodhead Publishing Limited, 2010
Ishige okamurae Sargassum fullvelum Sargassum horneri Sargassum coreanum Sargassum thunbergii Scytosipon lomentaria Undaria pinnatifida
Terrestrial plants Buckwheat (Fagopyrum esculentum) Caroleo (Olea europaea)
Enzymes
Extraction yield (total phenolics)
Antioxidant activity
Reference
Viscozyme Alcalase 2.4L FG Alcalase and Ultraflo
1352 mg GAE g–1 1095 mg GAE g–1
DPPH; HO˙; O2˙; H2O2
Heo et al., 2005
DPPH; H2O2
Ultraflo L Alcalase 2.4L FG Ultraflo L Alcalase 2.4L FG Ultraflo L Alcalase 2.4L FG Viscozyme Protamex Viscozyme Alcalase 2.4L FG Ultraflo L Alcalase 2.4L FG Alcalase, Flavourzyme, Promozyme, Viscozyme, Celluclast
275 mg GAE g–1 420 mg GAE g–1 313 mg GAE g–1 366 mg GAE g–1 384 mg GAE g–1 533 mg GAE g–1 1123 mg GAE g–1 996 mg GAE g–1 386 mg GAE g–1 416 mg GAE g–1 149 mg GAE g–1 207 mg GAE g–1
DPPH; HO˙; O2–; H2O2
Siriwardhana et al., 2008 Heo et al., 2005
DPPH; HO˙; O2–; H2O2
Heo et al., 2005
DPPH; HO˙; O2–; H2O2
Heo et al., 2005
DPPH; HO˙; O2–; H2O2
Heo et al., 2005
DPPH; HO˙; O2–; H2O2
Heo et al., 2005
DPPH; HO˙; O2–; H2O2
Heo et al., 2005
DPPH; HO–; O2–
Je et al., 2009
Alcalase
0.371 mg R g–1
DPPH; reducing power; inhibition of LA peroxidation
Tang et al., 2009
Bioliva
234.1 mg CA kg–1
Oxidative stability; peroxide value Ranalli et al., 2003b
532 Separation, extraction and concentration processes
Vegetal material (Latin name)
Leccino (Olea europaea) Leccino (Olea europaea) Lemon peel (Citrus limon cv. Meyer) Lemon peel (Citrus limon cv. Yen Ben) Mandarin peel (Citrus reticulara cv. Ellendale)
Bioliva
362.2 mg CA kg–1
Oxidative stability; peroxide value Ranalli et al., 2003a
Rapidase adex D
128 mg CA kg–1
Oxidative stability; peroxide value Ranalli et al., 2003a
Cytolase
170 mg CA kg–1
Oxidative stability; peroxide value Ranalli et al., 1999
Cytolase
79 mg CA kg–1
Oxidative stability; peroxide value Ranalli et al., 1999
Rapidase adex D
92 mg CA kg–1
Oxidative stability; peroxide value Ranalli et al., 2003
Cellulase
220 mg F g–1
Huang et al., 2009
Pectinases and cellulases
89 mg CA g–1
HO˙; O2–; inhibition of lipid peroxidation; reducing power DPPH; oxidative stability
Cellulase MX Cellulase CL Kleerase AFP Bioliva
1.52 mg GAE g–1 1.30 mg GAE g–1 1.25 mg GAE g–1 121.2 mg CA kg–1
FRAP
Li et al., 2006
Cytolase
121 mg CA kg–1
Oxidative stability
Ranalli et al., 1999
Cellulase MX Cellulase CL Kleerase AFP Cellulase MX Cellulase CL Kleerase AFP Cellulase MX Cellulase CL Kleerase AFP
0.4429 mg GAE g–1 0.4895 mg GAE g–1 0.4375 mg GAE g–1 1.131 mg GAE g–1 0.99 mg GAE g–1 1.108 mg GAE g–1 1.39 mg GAE g–1 1.15 mg GAE g–1 1.19 mg GAE g–1
FRAP
Li et al., 2006
FRAP
Li et al., 2006
FRAP
Li et al., 2006
Ramadan et al., 2008
Oxidative stability; peroxide value Ranalli et al., 2003
Extraction of natural antioxidants from plant foods 533
© Woodhead Publishing Limited, 2010
Coratina (Olea europaea) Coratina (Olea europaea) Coratina (Olea europaea) Dritta (Olea europaea) Dritta (Olea europaea) Eucommia leaf (Folium eucommiae) Goldenberry (Physalis peruviana) Grapefruit peel (Citrus x paradisi)
© Woodhead Publishing Limited, 2010
Vegetal material (Latin name)
Enzymes
Extraction yield (total phenolics)
Antioxidant activity
Moringa concanensis seeds
1.283 mg T kg–1 1.236 mg T kg–1 1.176 mg T kg–1
Oxidative stability; peroxide value; Latif and Anwar, conjugated diene/triene 2009
Sunflower (Helianthus annuus)
Kemzyme Natuzyme Feedzyme Protex 7L Kemzyme Alcalase 2.4L Viscozyme L Natuzyme
Sweet orange peel (Citrus sinensis cv. Navel)
Cellulase MX Cellulase CL Kleerase AFP
0.80 mg GAE g–1 0.79 mg GAE g–1 0.90 mg GAE g–1
FRAP
Li et al., 2006
Macer8 Fj (M) Grindamyl pectinase (G) (M) + (G) Olivex and Celluclast
358 mg GAE L–1 237 mg GAE L–1 383 mg GAE L–1 5.2 g C kg–1
Conjugated diene
Landbo and Meyer, 2001
DPPH
Soto et al., 2008
Grindamyl pectinase (G) Celluclast (C) (G) + (C) Ultrazym 100G and Cellubrix Ferulic acid esterase
3072 mg GAE L–1 2773 mg GAE L–1 3017 mg GAE L–1 0.25 g C g–1
Conjugated diene
Meyer et al., 1998
DPPH
Collao et al., 2007
6937 mg SAE g–1
DPPH; conjugated diene; hydroperoxide; hexanal; liposome model system
Vuorela et al., 2004
Residues Blackcurrant pomace (Ribes nigrum) Borage seeds (Borago officinalis) Grape seeds pomace (Vitis vinifera) Evening primrose seeds (Oenothera biennis) Rapeseed (Brassica rapa)
13 14 13 15 13
mg mg mg mg mg
GAE GAE GAE GAE GAE
g–1 g–1 g–1 g–1 g–1
Reference
DPPH; inhibition of LA; oxidation Latif and Anwar, 2009
C, catechin; CA, caffeic acid; F, flavonoid; GAE, gallic acid equivalents; R, rutin; SAE, sinapic acid equivalents; T, tocopherol.
534 Separation, extraction and concentration processes
Table 18.8 Continued
Extraction of natural antioxidants from plant foods 535 Extractable compounds (mg/100 g)
4.5
1
4.0 3.5
2
3
4
3.0 2.5 5
2.0 1.5
7
1.0
9
0.5 0
6
8
10 0
2
4 6 Enzyme concentration (%)
8
10
Fig. 18.2 Effect of enzyme concentration on the yield of extractable compounds. 1: Bilberry (Biopectinase CCN) (Puupponen-Piniä et al., 2008). 2: Bilberry (Pectinex BE 3L) (Puupponen-Piniä et al., 2008). 3: Blackcurrant (Biopectinase CCN) (PuupponenPiniä et al., 2008). 4: Blackcurrant (Pectinex BE 3L) (Puupponen-Piniä et al., 2008). 5: Pigeonpea (Fu et al., 2008); lutein (mg/100 g) ¥ 10–1; enzyme concentration ¥ 10–2. 6: Grape pomace (Maier et al., 2008); enzyme concentration ¥ 10–1; 7: Apple pomace (Pinelo et al., 2008). 8: Pigeonpea (Fu et al., 2008); apigenin ¥ 10–1; enzyme concentration ¥ 10–2. 9: Meyer lemon (Li et al., 2006). 10: Yen Ben lemon (Li et al., 2006).
Ecklonia cava showed antiproliferative action (Athukorala et al., 2006). The phlorotannin derivatives of this alga exhibited inhibitory effect on human immunodeficiency virus type 1 reverse transcriptase and protease (Ahn et al., 2004). Enzymatic extracts of Stellaria dichotoma inhibited the in vitro hydroxyl radical-induced DNA damage (Lim et al., 2008). A peptide with ACE-inhibitory activity was isolated from the pepsin hydrolyzate of Chlorella vulgaris waste (Sheih et al., 2009), whereas some glucans (hexaose, schizophyllan, and laminarin) showed inhibition against AAPH (Tsiapali et al., 2001). The major antioxidants of fruits and juices are phenolic compounds. Phenols may be found in cell walls, bound to polysaccharides by hydrophobic interactions and hydrogen bonds, linked to the protein matrix of vacuolar inclusions, confined inside the cellular vacuoles, and near or associated with the cell nucleus (Pinelo et al., 2006). Plant cell walls are a complex network of cellulose, hemicelluloses (mainly of xyloglucans, mannans, xylans, and arabinogalactans), pectins and lignin (a hydrophobic polymer derived from p-coumaryl, coniferyl, and sinapyl alcohols). Degradation of the cell-wall polysaccharide structure is a fundamental step in the release of phenols linked to the cell wall or associated with cell vacuoles. Ferulic and p-coumaric acids, the major lignin monomers that link hemicellulose sugars and lignin, are potent in vitro antioxidants (Meyer et al., 1998; Nardini et al., 1995). Enzymatic degradation of cell-wall polysaccharides is assumed to increase © Woodhead Publishing Limited, 2010
536 Separation, extraction and concentration processes the overall substrate porosity, facilitating solvent penetration and extraction efficiency, but lignin and tannins may block the enzyme action. The use of enzymes that degrade plant cell walls to enhance the extraction of phenolic antioxidant compounds has been reported for food products and by-products. Although no interactions between proteases and pectinases was confirmed for blackcurrant juice production, degradation of the protein network enhanced both substrate porosity and phenol recovery (Landbo and Meyer, 2001). In grape skins, cellulases, hemicellulases, pectinases, and other hydrolytic enzymes can be used (Pinelo et al., 2006). Despite the high cellulose content of the grape cell wall, cellulolytic activity did not influence the extraction yield of phenols from pomace (Meyer et al., 1998), but pectinases, cellulases and proteases increased the extraction yields from apple skin (Pinelo et al., 2008). Preferential enzymecatalyzed solubilization of pectin and cellulose from bergamot peel caused the selective release of glucose and galacturonic acid (Mandalari et al., 2006). Glucanases and pectinases were used for processing citrus peels (Li et al., 2006). Neutral protease, papain and alkaline protease have been assayed on the fruit of Physalis alkekengi (Ge et al., 2009), whereas culture broths of Aspergillus niger enriched in cinnamoyl esterases were used to release hydroxytyrosol from olive oil by-products (Bouzid et al., 2005). The combined effect of the major operational variables (including particle size, time, temperature, and enzyme to substrate ratio) is evaluated using experimental designs (Meyer et al., 1998). The particle size is relevant for degrading plant cell wall carbohydrates, as well as for solvent penetration and solute recovery. A favorable effect of the reduction of particle size on the extraction of phenolics was reported for spent coffee grounds (Pinelo et al., 2007), blackcurrant pomace (Landbo and Meyer, 2001), apple skin (Pinelo et al., 2008), and citrus peels (Li et al., 2006). Reduced particle size favored the degradation of polysaccharides by pectinase without affecting the recovery of phenols (Meyer et al., 1998). Although a beneficial effect of high enzyme concentration (5%) was reported for citrus peel, the optimal value for this material was 1–3% (Li et al., 2006). Increased enzyme dosage enabled higher juice yields and phenol concentrations from blackcurrant mash. However, the non-enzyme-treated juices exhibited higher antioxidant activity than the enzyme-treated ones, presumably because of differences in their phenolic profiles (Bagger-Jørgensen and Meyer, 2004). The effects of enzymatic treatments were different for various citrus fruits (lemon, mandarin, grapefruit and orange) (Li et al., 2006), and various effects of temperature, particle size reduction and ethanol concentration were reported for the extraction of individual compounds from apple skins (Pinelo et al., 2008). Optimal temperatures of 50–60 °C were reported for citrus peel (Li et al., 2006), apple skins (Pinelo et al., 2008), and elderberry (Landbo et al., 2007). Prolonged incubation time (24 h) with pectinolytic and cellulolytic enzymes was necessary to solubilize carbohydrates and flavonoids from bergamot peel (Mandalari et al., 2006), but a negative effect of long reaction
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Extraction of natural antioxidants from plant foods 537 times was observed for grape owing to degradation of phenolic compounds (Meyer et al., 1998). Enzyme-assisted aqueous of citrus peels required an incubation period of 3 h, compared with 6 h for water extraction without enzymes (Li et al., 2006). The amount of quercetin in enzyme-treated juices increased with incubation time, but the quercetin content did not involve a higher antioxidant activity (Sun et al., 2005). The yields of soluble carbohydrates were higher in enzyme-treated samples than in untreated ones. A positive linear correlation between extent of cell wall degradation and phenol extraction was observed for grape pomace (Meyer et al., 1998), elderberry (Bagger-Jørgensen and Meyer, 2004), and blackcurrants (Bagger-Jørgensen and Meyer, 2004; Landbo and Meyer, 2001), but no significant differences were found in carbohydrate degradation, distribution of phenolics and release of anthocyanins and total phenols (Landbo et al., 2007). Similarly, pectinolysis affected both juice yield and contents of phenolic acids, flavonoids and anthocyanins (Kaack et al., 2008). Even though the amount of phenols extracted from the pomace is enhanced, specific compounds may decrease: for example, enzyme-mediated anthocyanin degradation may occur owing to enzymatic hydrolysis of glycosylated anthocyanins by polyphenol oxidases. Pectinase preparations produced by Aspergillus niger, containing b-glucosidase, b-galactosidase, and a-l-arabinosidase, affect anthocyanin pigments and color quality (Landbo and Meyer, 2001). Despite b-galactosidase and b-glucosidase causing color degradation, benefits for extraction of potent antioxidants from grape pomace were reported (Meyer et al., 1998). Enzymatic release of esterified hydroxycinnamic acids can be achieved by synergistic action of cell-wall-degrading enzymes (esterases and xylanases). Commercial pectinases facilitate the release of ferulic acid and other phenolic acids from ground rye grain (Andreasen et al., 1999). A cold pressing process was used for defatting borago before extracting phenolics with conventional solvents. Enzymatic pretreatment of seeds enhanced the extraction of phenolic compounds with DPPH radical scavenging capacity from the meal; the tocopherol content of the resulting oil was maintained by cold pressing (Soto et al., 2008). Treatments with cell-wall-degrading enzymes were proposed for releasing ferulic acid from agro-industrial byproducts such as sugar beet pulp or maize bran (Bouzid et al., 2005). Other less studied materials include microbial sources. Saccharomyces cerevisiae b-glucan, isolated by hot water and enzyme treatment, exhibited antioxidant activity (Jaehrig et al., 2007). Protease treatments can be included within the enzyme-assisted extraction methods. Relatively short bioactive peptides (2–9 amino acids) can be obtained from sources such as wheat, corn, soybean and mushrooms (Dziuba et al., 2004). Commercial proteases have been used to process a number of substrates (Tang et al., 2009; Wang et al., 2007). Hydrolysis increased ACE inhibition and the radical-scavenging activity of protein hydrolyzates from potato isolates and by-products, possibly related to peptides and /or
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538 Separation, extraction and concentration processes free amino acids liberated during digestion (Pihlanto et al., 2008). Rapeseed protein hydrolyzates present DPPH radical-scavenging activity, provide protection against the peroxidation of lipids and/or fatty acids (Amarowicz and Shahidi, 1997; Peña-Ramos and Xiong, 2002; Sakanaka et al., 2004) and show reducing power and Fe2+ chelating ability (Moure et al., 2006; Tang et al., 2009). Esterified phenolic acids (p-coumaric acid, ferulic acid) present in sugarcane bagasse were released by alkaline hydrolysis (Ou et al., 2009). Ferulic acid (FA), a scavenger of free radicals approved in certain countries as a food additive to prevent lipid peroxidation (Srinivasan et al., 2007), can be extracted using enzymatic, alkaline or acidic extractions. 18.5.3 Pressurized-liquid extraction Pressurized-liquid extraction (PLE) is an efficient, innovative and environmentally clean technique, which is performed at high temperature and pressure to maintain the solvent in the liquid state. Compared with traditional extraction techniques, it is faster and requires lower amounts of solvent. Alternatively, it is known as accelerated solvent extraction (ASE), or high-pressure liquid extraction (HPE) (Adil et al., 2008). The structure, activity and properties of the extracts are usually unaffected in an oxygenfree and light-free environment, and micro-organisms and enzymes may be inactivated (Qadir et al., 2009). The influence of HPE on the antioxidant capacity of food products was recently reviewed by Oey et al. (2008). PLE was proposed for the extraction of a variety of natural products and wastes, including the production of anthocyanins from grape skins (Ju and Howard 2003), flavonoids from spinach (Howard and Pandjaitan, 2008) and processing of berry substrates (King et al., 2003). The use of non-toxic solvents (water and ethanol) offers environmental advantages. Ethanol–water mixtures gave high extraction yields from Phormidium species (Rodríguez-Meizoso et al., 2008), sour cherry pomace (Adil et al., 2008), and dried spinach (Howard and Pandjaitan, 2008); whereas methanol was used for extracting catechin and epicatechin from tea leaves and grape seeds (García-Marino et al., 2006; Piñeiro et al., 2004), and flavonoids from apple extracts (Alonso-Salces et al., 2001). The compounds recovered and their extraction efficiency differed with the solvent, providing different properties to the extracts, including color and odor compounds (Howard and Pandjaitan, 2008). The amount of carotenoids extracted from Dunaliella salina with hexane was more than seven times higher than the value extracted with ethanol, but the antioxidant activity was only twice as high owing to structural differences (Herrero et al., 2006). Combinations of pressure, temperature, solid/solvent ratio and extraction time can be addressed by experimental designs (Adil et al., 2008; Herrero et al., 2006). Temperature, vapor pressure of the most volatile compounds and improved mass transfer have been reported as the major variables involved in
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Extraction of natural antioxidants from plant foods 539 PLE regardless the solvent used (Rodríguez-Meizoso et al., 2008). Temperature was the most influential factor in the extraction of antioxidant compounds from Dunaliella salina (Herrero et al., 2006). The optimal temperature for water and ethanolic extraction of Spinacia oleracea flavonoids was 50–130 °C for water, and 50–150 °C for ethanol (Howard and Pandjaitan, 2008), compared with 100–130 °C for PLE processing of grape by-products (Ju and Howard 2003; Piñeiro et al., 2004) and berries (King et al., 2003). The optimal temperature depends on the solvent considered. Optimal antioxidant activity of Phormidium species in hexane extracts was obtained operating at 50 °C, whereas in ethanol and water extracts it was observed at 150 and 200 °C, respectively (Rodríguez-Meizoso et al., 2008). Thermal degradation of apple polyphenols was observed at extraction temperatures higher than 60 °C in methanol, whereas activity decreased at 110–190 °C in water and ethanol (Alonso-Salces et al., 2001). Thermal degradation of flavonoids was reported for extraction temperatures >130 °C (Howard and Pandjaitan, 2008). The reduction in antioxidant activity observed at prolonged extraction times during PLE of Dunaliella salina with hexane and ethanol suggested some kind of carotenoid degradation (Herrero et al., 2006). HPE extraction can give a moderately high extraction efficiency of bioactive compounds in a short time, but treatments for prolonged periods could result in degradation (Qadir et al., 2009). The development of brown, highly polar, odoriferous compounds with high antioxidant capacity at high extraction temperatures suggested the participation of Maillard reactions in ethanolic (Howard and Pandjaitan, 2008) and in water extracts (Rodríguez-Meizoso et al., 2008). Extracts of propolis obtained by high hydrostatic pressure extraction in 1 min presented similar properties to the ethanolic extracts obtained by leaching at room temperature for a few days (Jun, 2006). No degradation of compounds present in ethanolic extracts occurred after exposure to light and air at room temperature for 2 days (Herrero et al., 2006). Other reported biological activities include cytotoxic effects on the growth of human promyelocytic leukemia cells (HL-60) of sweet potato extracts (Rabah et al., 2005), toxicity on preformed monolayers of Vero cells (African green monkey kidney cell line) of Phormidium extracts from PLE (Rodríguez-Meizoso et al., 2008). Antimicrobial activity of Phormidium sp. extracts has been ascribed to terpenes (i.e., b-ionone, neophytadiene) and fatty acids (i.e., palmitoleic and linoleic acids) (Rodríguez-Meizoso et al., 2008). Subcritical-water extraction (SWE) is an environmentally friendly technology for the selective extraction of bioactive compounds from plant materials with advantages derived from its simplicity, reduced extraction time, high quality of the extract, low cost of the extracting solvent and possibility of scaling-up (Rodríguez-Meizoso et al., 2006). Extract with water under pressure (between 100 °C and the critical temperature of 374 °C) is also known as superheated water extraction, subcritical water extraction, high-temperature water extraction, pressurized hot water extraction, or hot
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540 Separation, extraction and concentration processes water extraction. Most studies have been performed in batch mode on a small scale, but continuous flow equipment has also been described (Budrat and Shotipruk, 2008). The feedstocks studied include fruits or vegetables and their processing wastes, seeds, herbs and algae (Table 18.9). SWE was more efficient than conventional solvents for extracting flavanols from grape seeds (García-Marino et al., 2006) or antioxidant phenolics from bitter melon (Budrat and Shotipruk, 2008), and was sucessfully employed to concentrate and isolate antioxidant compounds from oregano leaves (Rodríguez-Meizoso et al., 2006). Temperature has a marked effect on yield and selectivity. The dielectric constant of water decreases with temperature, enabling the extraction of compounds having different polarities at different temperatures. High temperature also enhances diffusivity, facilitating the transport of solutes from the solid matrix, and a compromise to avoid thermal degradation must be reached. The release of hydroxycinnamates from cell walls is favored by temperature, but anthocyanins undergo degradation reactions. Lignin decomposes in sub- and supercritical water to give phenol, which are degraded at high temperatures (Budrat and Shotipruk, 2008; Garrote et al., 2003). Sequential extraction with stepwise pressure increase was proposed for processing black tea leaves (Chambers et al., 1984) as well as to recover quercetin glycosides from onion waste (Turner et al., 2006), and catechins and proanthocyanidins from winery by-products (García-Marino et al., 2006). The major variable is temperature, because at 150 °C or above, substrates are hydrolyzed, and the solvent polarity decreases. Optimization of operational conditions has to be carried out for each material. The yield of total phenolics from sweet potato and bitter melon increased with temperature (Budrat and Shotipruk, 2008; Rabah et al., 2005). Although higher temperatures resulted in higher amounts of extracted phenolics, their antioxidant activity was higher operating at lower temperatures (Budrat and Shotipruk, 2008). Carnosic acid, the most potent antioxidant in rosemary, was preferentially extracted at 200 °C (Rodríguez-Meizoso et al., 2006). Extraction yields of oregano leaves increased with temperature: at the lowest temperature considered, the more polar compounds (flavanones and dihydroflavonol structures such as dihydroquercetin, eriodictyol and dihydrokaempferol) were preferentially extracted, whereas less polar compounds could also be extracted at 200 °C (Rodríguez-Meizoso et al., 2006). The phenolic content was a maximum at 25 °C and a minimum at 200 °C, but the highest antioxidant activities were achieved at 150–200 °C. Boiling ground coffee beans in water at 110 °C under elevated pressure gave an efficient extraction of antioxidants (Budryn and Nebesny, 2008). The total phenolic contents and the antioxidant activities of HPE extracts from longan fruit were higher than those obtained with conventional solvent extract (Prasad et al., 2009), whereas enhanced extraction of onion flavonols was observed at 100–400 MPa (Roldán-Marín et al., 2009). The conditions of SWE influence the degree of polymerization and structure
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Table 18.9 Examples of extraction of plant food antioxidants using high pressure systems Raw material
Operational conditions
Elderberry leaves (Sambucus nigra L.) Elderberry berries (Sambucus nigra L.) Elderberry flowers (Sambucus nigra L.) Phormidium species
Red grape skin (Vitis vinifera)
Antioxidant activity Reference
0.32 mg GAE g–1
DPPH
Bonoli et al., 2004
Fiber: 14.99 g GAE kg–1 Flour: 20.93 g GAE kg–1
DPPH
Papagiannopoulos et al., 2004
28.91 mg GAE g–1
DPPH ABTS FRAP DPPH bC bleaching
Güçlü-Üstündağ and Mazza, 2009
DPPH bC bleaching DPPH bC bleaching TEAC Antimicrobial activity
Dawidowicz et al., 2006 Dawidowicz et al., 2006 Rodríguez-Meizoso et al., 2008
ORAC
Ju and Howard, 2003
1 g; ECV(22); 17 (g F/100 g ¥ 10–2) SE: E:W (80:20); -; 10 min; 60 bar; 200 ºC; PLE: FV( 60% ECV); P(120) 20.18 (g F/100 g ¥ 10–2) 214.2 (g F/100 g ¥ 10–2) 1 g sample H: 20 min; 200 ºC; E: 20 min; 200 ºC; W: 20 min; 200 ºC; 0.2 g + 30 g sea sand; ECV(22); SE: acidified W; 3 cycles; 5 min; 10.1 MPa; 50 ºC PLE; FV (50% ECV); P(90)
6.54% 40.91% 22.81% 179.9 mg GAE g–1
Dawidowicz et al., 2006
Extraction of natural antioxidants from plant foods 541
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Pressurized liquid extraction (PLE) Barley flour 2 g + 4 g Hydromatrix; ECV(33); (Hordeum vulgare L.) SE: E:W (4:1); 2 cycles; 5 min, 20 MPa 60 °C PLE: FV (60% ECV); P(60) Carob fruits 1 g + 2 g diatomaceous earth; (Ceratonia siliqua L.) SE: A:W (1:1); 2 cycles; 5 min; -; 60 °C; PLE: 50% FV; Cow Cockle seeds 2 g; (Saponaria vaccaria) W; 3 h; 125–175 °C; F(2)
Solubles yield or TEP*
Raw material
Operational conditions
Spinach (Spinacia oleracea)
0.5 g + 30 g Ottawa sand 42 mmol GAE kg–1 SE: E:W (7:3); 3 cycles; 5 min; 13.8 MPa; -; PLE: R (70% ECV) 2.5 g; ECV(11) PE; 9 min; 170 ºC; 77.0 mg extract H; 15 min; 170 ºC; 95.5 mg extract E; 9 min; 170 ºC; 270.1 mg extract W; 15 min; 170 ºC; 290.0 mg extract 2.5 g; ECV(11) PE; 9 min; 170 ºC; 2.94% H; 9 min; 170 ºC; 4.28% E; 9 min; 170 ºC; 19.70%
Spirulina plantensis © Woodhead Publishing Limited, 2010
Spirulina plantensis
Solubles yield or TEP*
Antioxidant activity Reference ORAC
DPPH
DPPH
Howard and Pandjaitan, 2008 Herrero et al., 2004
Jaime et al., 2005
High pressure extraction (HPE) Korean barberry bark (Berberis koreana)
1 g mL–1; 15 min; 500 MPa; RT
50% E; 0.02 g mL–1; 30 min; 500 MPa; 50 ºC E; 0.06–0.07 g mL–1; 25 min; 176–193 MPa; 60 °C High pressurized fluid extraction (HPFE) Brazilian propolis RSL 1:10; 20% E; 40 g surfactant; 50 psig; 393 K Longan fruit pericarp (Arillus longan) Sour cherry pomace
11.04% (317.35 mg GAE g–1)
DPPH Inhibition of xanthine oxidase
Qadir et al., 2009
30% (23 mg GAE g–1)
Prasad et al., 2009
3.80 mg GAE g–1
DPPH O2– DPPH
Adil et al., 2008
25.2%
DPPH
Chen et al., 2007
542 Separation, extraction and concentration processes
Table 18.9 Continued
Taiwan Red Pine (Pinus taiwanensis)
70% E; 180 min; 689 kPa; 343 K; F (10)
26.3 mg g–1
Taiwan White Pine (Pinus morrisonicola)
70% E; 180 min; 689 kPa; 343 K; F (10)
39.8 mg g–1
Propolis crude
RSL 1:35, E 75%; 1 min; 500 MPa; RT
Accelerated solvent extraction (ASE) Spirulina plantesis 2.5 g; ECV (11); 15 min; -; 170 °C
–
DPPH
Rieger et al., 2008
–
DPPH
Rieger et al., 2008
317 mmol GAE g–1 dm
ABTS
Corrales et al., 2008
–
DPPH
Rieger et al., 2008
TPA: 0.20, TF: 7.10, TOD: DPPH ORAC 63.5 mg GAE g–1
Súarez et al., 2009
290 mg GAE g–1
bC bleaching DPPH
Xi, 2006
6.43 g GAE g–1
bC bleaching DPPH
Xi and Shouqin, 2007
H: 4.3%; PE: 4.01% E: 17.14%; W: 10.12%
DPPH
Herrero et al., 2005
Extraction of natural antioxidants from plant foods 543
© Woodhead Publishing Limited, 2010
High hydrostatic pressure (HPP) Bilberry ECV(11); (Vaccinium myrtillus) SE: 3 cycle, 7 min, 60 bar, 80 ºC; PLE: FV (100%), P (100) Elderberry ECV (11); (Sambucus nigra) SE: 3 cycle; 5 min; 68.9 bar; 60 °C; PLE: FV (100%), P (60) Grape skins RSL 1:4.5; E:W (1:1); 1 h; 600 MPa; (Dornfelder V. vinifera) 70° C; Heather ECV (11); (Calluna vulgaris) SE: 3 cycle, 5 min; 60 bar; 60 ºC PLE: FV (100%), P (60) Olive cake 30 g OC; ECV (100) E:W (80:20), FV (60% ECV), 2 cycles; 5 min Propolis extract RSL 1:35, E 75%; 1 min; 500 MPa; RT
DPPH, Fe Lin et al., 2009 chelating, reducing ability, O2–, NO DPPH, Fe Lin et al., 2009 chelating, reducing ability, O2– NO
Raw material
Operational conditions
© Woodhead Publishing Limited, 2010
Solubles yield or TEP*
Antioxidant activity Reference
Hot water extraction (HWE) Canola meal RSL 1:16.7; W; 30 min; -; 80 °C (Brassica napus)
0.15 g g–1; 7.83 g SA g–1
Citrus peels
RSL 1:20; bddW; 30 min; -; 100 ºC
~32.5 mg GAE g–1 dw
Formosa koa leaves (Acacia confusa)
bdd H2O; 4 h; -; -
7.3 g
Green tea
RSL 1:100; dW; 5–30 min; -; 80 ºC
Herbal tea (Anagallis arvensis) Japanese horse chestnut (Aesculus turbinata ) Kaffir lime fruit peel (Citrus hystrix) Lettuce extract (Lactuca sativa) Mallotus japonicus leaves
RSL 1:166.7; bW; 5 min;- ; -
TP: 2050 mg GAE L–1 TF: 1700 mg GAE L–1 1.47 mmol Trolox g–1
RSL 1:100; bW; 2 h
1560 mg (280 mg EE)
DPPH TEAC Reducing power bC bleaching DPPH FRAP DPPH NBT Reducing power FRAP ABTS CUPRAC ABTS DPPH
0.2 g; ECV (10.8); 15 min; 200 ºC
23.7 mg GAE mL–1
DPPH
RSL 1:2; W; 80 ºC; 10 min
–
RSL 1:50 mL; bdW; 20 min
7.5 mg GAE mL–1
Morinda citrifolia roots
0.5g, ECV(10); 4 MPa; 2 h; 170 °C F (5)
92.55%
Peroxidation of liposomes DPPH O2– HO· DPPH
Hassas-Roudsari et al., 2009 Xu et al., 2008 Tung et al., 2009 Rusak et al., 2008 Apak et al., 2006 Ogawa et al., 2008 Khuwijitjaru et al., 2008 Altunkaya et al., 2008 Tabata et al., 2009 Pongnaravane et al., 2006
544 Separation, extraction and concentration processes
Table 18.9 Continued
Samor thai fruits (Terminalia chebula Retz.) Sasa palmata leaf
15.96 mg TP g–1 dw
ABTS
Rangsriwong et al., 2009
SE: 50 g; steam; 0.5–20 min; 180–260 °C (1.0–4.9 MPa) HWE: RSL 1:100; dW; 2 h; 98 °C RSL 1:100; dW; 2 h; 98 ºC
128.24 mg GAE g–1
DPPH
Kurosumi et al., 2007
12.08 mg GAE g–1
DPPH
Kurosumi et al., 2007 Spigno and De Faveri, 2009 Rusak et al., 2008
Tea
RSL 1:100; bW; 30–210 s
White tea
RSL 1:100; dW; 5–30 min; 80 ºC
Hot pressurized water extraction (HPWE) Boldo RSL 1:10; W; 3 h; 110 ºC (Peumus boldus M.) Sage 0.15 g + sea sand; 60 min; 100 kg cm–2; 100 ºC; F(1) (Salvia officinalis) Yam (Dioscorea alata)
LSR: 20 kg kg–1; ECV(1000); W; 3 h; 1.34 MPa; 120 ºC; F(10)
Subcritical water extraction (SWE) Bitter melon 1 g; ECV (10); (Momordica charantia) 3 mL min–1; 60 min, 200 °C, 10 MPa Canola meal 30 min; 6.89 MPa; 160 °C; F(1)
ABTS TP: 1300 mg GAE L–1 TF: 900 mg GAE L–1
FRAP ABTS
51.4% 5.2 mg B kg–1
ABTS DPPH
Del Valle et al., 2005 Ollanketo et al., 2002
2.143 g kg–1
DPPH
Chen et al., 2004
53 mg GAE g–1
ABTS
0.45 g g–1 15.42 g SA g–1
TEAC bC bleaching DPPH Reducing power
Budrat and Shotipruk, 2008 Hassas-Roudsari et al., 2009
Extraction of natural antioxidants from plant foods 545
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RSL 1:150; 100 °C; 120 min
© Woodhead Publishing Limited, 2010
Raw material
Operational conditions
Solubles yield or TEP*
Antioxidant activity Reference
Eucalyptus leaves (Eucalyptus globulus) Licorice roots (Glycyrrhiza glabra) Oregano (Origanum vulgare)
60 g; ECV (300) 200 mL W; 30 min; 4.91 MPa; 200 ºC 0.1 g; ECV (18.57) 10 mL W; 60 min; 0.002–5 MPa; 300 °C 0.75 g; ECV (11) SE:W; 15–30 min; 1500 psi; 25–200 °C; SWE: R (70% ECV) 0.40 g + 30 g sea sand; ECV (22) SE: 1 cycle; 40 s; – 110 ºC; SWE: R (70% ECV), P (90) 1.0 g; ECV (10) 37.5 min; 4 MPa; 180 °C; F (4) RSL 1:50; 10–60 min; 0.47–4.76 bar; 80–150 °C
4.8 mg g–1
Peroxynitrite
1521.3 ppm GAE 54% (0.149 g GAE g–1)
Radical scavenging Baek et al., 2008 Reducing power DPPH Rodriguez-Meizoso et al., 2006
52.3 mg TP g–1
ORAC
Ju and Howard, 2005
21.43 mg TP g–1
ABTS
0.73 mg TC g–1 dm 5.1 mg TF g–1 dm
ABTS
Rangsriwong et al., 2009 Prommuak et al., 2008
Red grape skin Samor thai fruits (Terminalia chebula) Thai silk waste *
Kulkarni et al., 2008
Total extractable polyphenols. B, boldine; bdW, boiled distilled water; bddW, boiled double distilled water; bW, boiled water; bC, b-carotene bleaching; dm, dry matter; dW, distilled water; E, ethanol; ECV, extraction cell volume; EE, epicatechin equivalent; F, flux (mL/min); FV, flush volume; GAE, gallic acid equivalent; H, hexane; LSR, liquid–solid ratio; P, purge (s); PE, petroleum ether; PLE, pressurized liquid extraction; pm, plant material; R, rising; RT, room temperature; SE, static extraction; TC, total carotenoids; TF, total flavonoids; TP, total phenolics; W, water.
546 Separation, extraction and concentration processes
Table 18.9 Continued
Extraction of natural antioxidants from plant foods 547 of molecules, as reported for procyanidin extraction (García-Marino et al., 2006). Enhanced recoveries of flavanol dimers and trimers, and gallic acid was obtained in greater quantities by a single extraction at 150 °C, in which gallic acid accounted for 61% of total phenolics. Higher temperatures favored the extraction of oligomeric fractions. When the galloylated moieties were situated in the terminal sub-unit, sequential extraction (50–100 °C) was preferable, whereas a single extraction at 150 °C was better for the rest of the galloylated derivatives (García-Marino et al., 2006). Treatment of lignocellulosic substrates with water or steam at 160–240 °C (autohydrolysis processing) results in depolymerization of hemicelluloses and in breakage of lignin-carbohydrate bonds, leading to solubilization of lignin fragments of low MW. The aqueous treatments of lignocellulosics for both the hydrolytic degradation of hemicelluloses and the solubilization of antioxidant compounds have been reviewed (Garrote et al., 2004; Meireles, 2009). Other related technologies, based on the utilization of chemicals (for example, solvents, oxygen or acids) added to water, were proposed to provide higher extraction yields. Steam treatments were efficient in releasing hydroxytyrosol and 3,4-dihydroxyphenylglycol (DHPG), from the semisolid waste of olive oil extraction systems (alperujo) (Fernández-Bolaños et al., 2002; Rodríguez et al., 2009). The amount of solubilized hydroxytyrosol increased with temperature and time, reaching 1.4−1.7 g/100 g of dry alperujo (FernándezBolaños et al., 2002). 18.5.4 Supercritical-fluid extraction Supercritical-fluid extraction (SFE), based on the utilization of a fluid under supercritical conditions, is suitable for the extraction and purification of natural compounds with bioactive and antioxidant properties. The interest in SFE is promoted by increasingly restrictive environmental, toxicological, and health regulations. Carbon dioxide shows favorable properties, including environmental safety, high availability at low cost, high purity, and suitability for extracting heat-labile, natural compounds with low volatility and polarity. The physicochemical properties of supercritical CO2 (high diffusivity, low viscosity and low surface tension compared with conventional solvents) facilitate mass transfer, enabling an environmentally friendly operation. The properties of the final product are improved by high selectivity, low processing temperatures, absence of light and oxygen, and absence of solvent in extracts. The extracts are regarded as natural, and do not require additional sterilization. The major disadvantages of supercritical CO2 are the high critical pressure (requiring expensive equipment), and the poor solvent power derived from its low polarity. The use of cosolvents enhances the solubility of some compounds and may increase the extraction selectivity, allowing operation at a lower pressure. Representative antioxidant compounds extracted by SFE are summarized
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548 Separation, extraction and concentration processes in Table 18.10. SFE successfully recovered phenolic compounds from liquid solutions (Gamse, 2004), and countercurrent contact has been employed for processing propolis (Wang et al., 2004), B. kaoi (Wang et al., 2005) and orange juice (Señoráns et al., 2001). Studies on the SFE of bioactive compounds with antioxidant activity are available (Del Valle et al., 2005; Díaz-Reinoso et al., 2006; Herrero et al., 2006; Meireles, 2003; Mukhopadhyay, 2000; Reverchon, 1997). The most influential variables in SFE of compounds with antioxidant properties from natural sources are summarized here. The drying technique (Mechanical and thermal conditioning) affects both the structure of the solid matrix and the thermal stability of the active compounds. Mass transfer is facilitated by a higher degree of grinding, but the performance of fixed beds during leaching could be limited by too fine particles. Pressure and temperature are the variables that determine the solubility equilibrium, an understanding of which is essential for process design. Results are mostly available for a single solute in the supercritical solvent, but the presence of other solutes may affect the equilibrium, and the natural products are a complex mixture of compounds. Studies dealing with the solubility of solid mixtures of phenolic compounds in supercritical CO2 have been reported (Del Valle et al., 1998; Díaz-Reinoso et al., 2006; Fornari et al., 2005; Lucien and Foster, 2000; Reverchon, 1997). Solvent density increases with pressure, but beyond a certain threshold, the increased solvent viscosity reduces the diffusion coefficients. When the target compounds are lipophilic, the operational conditions may be milder than those required for phenolic antioxidants. Highly labile compounds require mild temperatures (below 50 °C) to avoid alteration. A stepwise increase in the extraction pressure allows selective extraction of different compounds. The presence of a modifier increases the solvent density resulting in a higher interaction of solutes with the solvent, causing alterations and swelling of the vegetal matrix. A suitable cosolvent may improve the extraction yield and selectivity. However, cosolvents of hazardous nature or resulting in decreased selectivity should be avoided. The common modifiers used in SFE are alcohols, which induce dipole/dipole interactions and hydrogen bonding with polar functional groups. Ethanol has been employed to increase the solubility of ginsenoids (Wang et al., 2001), antioxidants from tamarind seed coat (Luengthanaphol et al., 2004), Eucalyptus (El-Ghorab et al., 2003) and olive leaves (Tabera et al., 2004), lipophilic compounds from marigold (Baumann et al., 2004), and flavonoids and terpenoids from Ginko biloba (Yang et al., 2002). Methanol was used for the extraction of magnolol (Dean et al., 1998), soy isoflavones (Rostagno et al., 2002), phenolics from chamomille (Scalia et al., 1999), pistachio hulls (Goli et al., 2005) or aloe extracts (Hu et al., 2005). Isopropanol was used for the extraction of ginger oleoresin (Zancan et al., 2002). Water is the cosolvent employed in the industrial extraction of vanilla. The extracts resulting from the unselective extraction of compounds at high pressure can be fractionated in multiple-stage separators. The pressure and
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Table 18.10 Examples of extraction of plant food antioxidants using SFE Extraction yield (total extract, TE and active compounds)
Antioxidant activity1
Reference
Aloe vera leaf skin (Aloe barbadensis) Baical skullcap root (Scutellaria baicalensis) Boldo (Peumus boldus) Boldo (P. boldus) Coriander seeds (Coriandrum sativum) Eucalyptus leaf oil (Eucalyptus camaldulensis) Eucalyptus leaves (E. camaldulensis)
TF = 1.50 TF = 0.27
DPPH DPPH
TE = 2.9; B (0.0031)e TE = 4.9; B (7.4)e TE = 1.92 MTp H (0.242); STp (0.0825) pCl (1.16); oHC (0.17); MEP (0.34); Thy (0.98)a TE = 16.6
TEAC ABTS DPPH Inhibition of LA oxidation Inhibition of LA oxidation
Hu et al., 2005 Bimbato and Tamanini, 2003 Del Valle et al., 2005 Del Valle et al., 2005 Yépez et al., 2002 Fadel et al., 1999
Ginger (Zingiber officinale)
Monoterpenes, sesquiterpenes and gingerols
b-C, LA
Leal et al., 2003
Ginger (Z. officinale)
TE = 2.31
b-C, LA
Zancan et al., 2002
Helichrysum dried flower heads (Helichrysum italicum) Hop (Humulus lupulus) Lemon balm (Melissa officinalis subs. officinalis) (M. officinalis subs. inodora) Marjoram (Origanum vulgare) Hungarian marjoram (O. majorana) Egyptian marjoram Nigella seeds (Nigella sativa)
Flavonoids; TE = 4.5–4.9
O2–;
Marongiu et al., 2003
TP = 0.52–3.79/TF = 0.2–0.92 TE = 1.9 TE = 0.7
LA oxidation LA oxidation
Lermusieau et al., 2001 Marongiu et al., 2004
TE = 3.2 TE = 3.76 TE = 5.39 T.O. (1.14)b; Tq (0.49); p-Cy (0.23)
Peroxide value Inhibition of oil oxidation
Uy et al., 1991 Vági et al., 2005
b-C, LA
Machmudah et al., 2005
DPPH; b-C, LA
El-Ghorab et al., 2003
Extraction of natural antioxidants from plant foods 549
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Vegetal material (Latin name)
© Woodhead Publishing Limited, 2010
Vegetal material (Latin name)
Extraction yield (total extract, TE and active compounds)
Antioxidant activity1
Reference
Black pepper oleoresin (Piper nigrum)
RER = 5.43 (Pp: 39.4)
Tipsrisukond et al., 1998
Rosemary (Rosmarinus officinalis) Rosemary (R. officinalis)
TE = 5.2 TE = 7.15
Hexanal formation; TBARS Peroxide value
Rosemary (R. officinalis) Rosemary (R. officinalis)
Carnosic acid solubility TE = 4.5; Cm (39.6); Vbn (20.3); t-C (10.3)a
Rosemary (R. officinalis)
Sage (Salvia officinalis) Sage (S. officinalis)
F1: Rsl (3.6); Gk (11.9); Cr (5.6); CrA (66.0); MCr(1.6)a F2: Rsl (8.0); Gk (0.13); Cr (6.4); CrA (60.7); MCr(4.6)a TE = 1.6; Cm (0.025 ); Cn (0.001); CrA (0.236); RA(0.065)c TE = 5.7 TE = 5.02
Sage (S. officinalis)
TE = 46.26; F1 (024)/F2 (27.50)/F3 (18.52)
Rosemary (R. officinalis)
Black sesame seed (Sesamum indicum) Black sesame seed (S. indicum) Summer savory (Satureja hortensis)
TE = 51.83 Lignans (sesamin, sesamolin) S1: Crv (32); Lnl (24); Myc (22); Myl (21)c S2: Crv (44); Lnl (12); Myl (11); g-Tpnn (20)c S3: Crv (55); g-Tpnn (38)c Tamarind seed coat (Tamarindus indica) 0.13 mg EC g–1
b-C, LA Peroxide value; TBARS
Uy et al., 1991 Dapkevicius et al., 1998 Ramírez et al., 2005 Leal et al., 2003
b-C, LA DPPH
Cavero et al., 2005
b-C, LA
Carvalho et al., 2005
Peroxide value
Uy et al., 1991 Dapkevicius et al., 1998
AAb-c PF
Daukšas et al., 2001
DPPH; Lipid oxidation DPPH; lipid oxidation Inhibition oil oxidation
Hu et al., 2004 Xu et al., 2005a Esquível et al., 1999
Peroxide value
Luengthanaphol et al., 2004
550 Separation, extraction and concentration processes
Table 18.10 Continued
Tamarind seed coat (T. indica)
Thyme leaves (T. vulgaris) Thyme (T. vulgaris)
TE = 4.92 TE = 0.7; Thy (0.17)d
Turmeric oil (Curcuma longa) Turmeric (C. longa)
TE = 5.5 (STp H)
Green tea (Camelia sinensis)
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1
Zg A (40.1); Eg A (18.2); arT (15.7); Zg (3.1)
Peroxide value
Tsuda et al., 1995
Lipid oxidation
Chang et al., 2000
IPDEBTA Peroxide value
Ko et al., 2002 Uy et al., 1991 Dapkevicius et al., 1998
b-C, LA Rancimat Rancimat AOP
a
b-C, LA
Simándi et al., 2001 Zeković et al., 2001 and 2003 Gopalan et al., 2000 Leal et al., 2003
Antioxidant activity reported in the referenced work or, if not indicated, other works referencing antioxidant activity of the plant extracts. Relative percentage of the normalized area detected by LC–MS or GC. b Total oil (essential oil). c Relative weight content in the extract; d mg (100 g)–1. e mg g–1. TE, total extraction yield (%); TF, total flavonoids extraction yield (%); TP, total phenolics extraction yield (%); RER, relative extraction rate. AAb-c = antioxidant activity during bleaching of b-carotene–linoleate solution, scale (0:Low–5:High). AOP = antioxidant potency with linoleic acid in an iron/ascorbate system. b-C, LA = absorbance of the b-carotene–linoleic acid system. Hexanal = percentage of reduction in the formation of hexanal respect to control. IP DEBTA = inhibition of linoleic acid peroxidation by the DEBTA (diethyl-2-thiobarbituric acid) method. PF = protection factor calculated as the ratio between the induction period of the sample with additive and the control sample during vegetable oil oxidation. TBARS = percentage of reduction in the formation of TBARS with respect to control. TEAC = Trolox equivalent antioxidant capacity. O 2– = superoxide radical scavenger capacity of superoxide dismutase. Simple phenolics, derivatives and flavonoids: B, boldine; Crv, carvacrol; DPA, 3,4-dihydroxyphenyl acetate; EgA, E-g-atlantone; EC, epicatechin; ECG, epicatechin gallate; EGCG, epigallocatechin gallate; GA, gallic acid; Gk, genkwanin; HDA, 2-hydroxy-3¢,4¢-dihydroxyacetophenone; MDB, methyl 3,4-dihydroxybenzoate; MEP, 2-methyl-6-ethylphenol; RA, rosmarinic acid; Thy, thymol. Terpenoids: t-C, trans-caryophylene; Cm, camphor; Cn, cineole; Cr, carnosol; CrA, carnosic acid; p-Cy, p-cymene; MCr, methyl carnosate; Lnl, linalool; MTp H, monoterpenes hydrocarbons; Myc, myrcene; Myl, myrtenol; Rsl, rosmanol; STp, sesquiterpenes; STp H, sesquiterpene hydrocarbons; ar-T, ar-turmerone; g-Tpnn, g-terpinene; Vbn, verbenone; ZgA, Z-g-atlantone; Zg, zingiberene. Other compounds: pCl, p-cymen-7-ol; oHC, O-hydroxycumine; Pp, piperine; Tq, thymoquinone. a
Extraction of natural antioxidants from plant foods 551
Terminalia catappa leaves, seeds Thyme (Thymus vulgaris) Thyme (T. vulgaris)
TE = 0.29; HDA (0.043); MDB (0.068); DPA (0.885); EC (0.164)d EGC (290); EGCG (510); ECG (105); EC (90); GA (7)e Tannin, flavonoid glycosides TE = 2.0 TE = 5.46
552 Separation, extraction and concentration processes temperature in the separators influence the fractionation of extracts. Series of two or three separation stages are used to recover the desirable components. Although the extraction yields are increased under harsh conditions, the selectivity is reduced, leading to extracts with poor antioxidant activity. These instances require optimization of the extraction stage (Gelmez et al., 2009) and/or separation steps. The yield can be similar to the use of a stepwise increase in the extraction pressure, but with a lower consumption of solvent. 18.5.5 Other novel extraction technologies The fundamentals and application of alternative techniques using solvents at low pressure for the extraction of bioactive compounds have been comprehensively reviewed (Meireles, 2009). Some of these techniques have been developed for analytical purposes, but they could be scaled up. Studies dealing with the antioxidant activity of extracts obtained with ultrasonic- and microwave-assisted technologies are summarized in Table 18.11. Ultrasound-assisted extraction (UAE) is a cheap, scalable technique allowing reduced extraction times, decreased extraction temperature, and increased extraction yields. Its effects are mainly attributed to cavitation forces upon the propagation of the acoustic waves. Ultrasounds result in physical, chemical, and mechanical effects promoting the release of soluble compounds from the plant body, enhancing mass transfer and facilitating the access of solvent to cell contents (Knorr et al., 2002; Ma et al., 2008). UAE was used to extract valuable compounds such as phenolic acids and flavanone glycosides (Ma et al., 2008). In UAE, both treatment time and (particularly) temperature have significant effects on the properties of extracts. However, UAE must be applied carefully to avoid degradation of susceptible solutes (Ma et al., 2008). Degradation of phenolics depends on their substitution pattern. High comparative stability has been reported for sinapic and vanillic acids, which have methoxyl-type substituents in their aromatic rings, whereas benzoic acids are more stable than cinnamic acids. In previously dried samples, ultrasound accelerated rehydration and swelling, without significant chemical degradation (Qadir et al., 2009). The ultrasonic power had a positive effect on the contents of phenolic acids from mandarin peels (Ma et al., 2008), tannins from myrobalan nuts (Sivakumar et al., 2008), and anthraquinones from Morinda citrifolia roots (Hemwimol et al., 2006). Differences in the effects of acoustic treatments can be ascribed to the utilization of different equipment, effects of the nature of the solid matrix (hardness, compactness, and solute distribution), and cavitation. Some of the general trends outlined above were confirmed in the extraction of active components from Hypericum perforatum L. (Smelcerovic et al., 2006). Other technologies enhancing mass transfer in plant tissues and permeability of cytoplasmatic membranes include microwave-assisted extraction (MAE), pulsed electric fields (PEF) and high hydrostatic pressure (HHP) (Corrales
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Table 18.11 Other extraction technologies applied to the recovery of antioxidants from food plants Operational conditions
Extraction yield
Antioxidant activity
Reference
Ultrasound Almond shells
LSR 1:10; 5% NaOH; RT; 10 min
TEP (%)10.1
DPPH
Bearberry leaves
LSR 1:25; W:E; 60 kHz; 25 ºC; 40 min
2.15 %
Cow cockle seeds (Saponaria vaccaria)
LSR 1:50; (W, 50% E, ACN); RT; 1 h
W: 0.78 g GAE g–1 dw 50%E: 1.58 g GAE g–1 dw ACN: 0.20 g GAE g–1 dw
Dornfelder skins (Vitis vinífera) Ginseng (Panax ginseng) Ginseng (P. ginseng) Indian mulberry (Morinda citrifolia) Indian mulberry roots (M. citrifolia) Mandarin peels (Citrus unshiu)
35 kHz; 70 °C; 1 h
360 mmol GAE g–1 dw
Inhibition of oil oxidation DPPH ABTS FRAP ABTS
Ebringerová et al., 2008 Gribova et al., 2008
Corrales et al., 2008
LSR 1:50; 70% E; 60 ºC; 40 min; 250 W; 42 kHz LSR 1:50; 70% E; 70 °C; 2 h; 42 kHz; 284 W RLS 100; E; 60 °C; 60 min; 15.7 W
TS: 3.89 %
DPPH
Chen et al., 2010
G: 13.7 mg g–1 TP: 7.1 mg g–1 62.23%
DPPH
Kim et al., 2007
DPPH
79.62%
DPPH –
Pistachio (Pistachia vera) Plum (Prunus domestica)
LSR 1:8; W; M; 45 min
TCE:1935.12 mg g–1 dw FG: 1374.35 mg g–1 dw W: 34.2 mg TAE g–1 dw M: 32.8 mg TAE g–1 dw TP:1.74–3.75 mg GAE g–1 TP:1.18–2.37 mg CE g–1
Hemwimol et al., 2006 Pongnaravane et al., 2006 Ma et al., 2008
Peroxide value
Goli et al., 2005
ABTS
Kim et al., 2003
LSR 1:10; E; 60 ºC; 2 h; 38.5 kHz; 270 W LSR 1:80; 80% M; 30 °C; 10 min; 8 W
LSR 1:10; 80% M; -; 20 min
Güçlü-Üstündağ and Mazza, 2009
Extraction of natural antioxidants from plant foods 553
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Raw material
© Woodhead Publishing Limited, 2010
Raw material
Operational conditions
Extraction yield
Antioxidant activity
Reference
Sage (Salvia officinalis) Sea-buckthorn (Hippophae rhamnoides) Sempervivum (Sempervivun marmoreum) Wheat bran (Triticum aestivum) Microwave
LSR 1:4; M; RT, 20 min
–
DPPH
LSR 1:10; E; 30 °C; 60 min
15.6 (0.39) mg GAE g–1
DPPH TEAC
Ollanketo et al., 2002 Sharma et al., 2008
LSR 1:10; dW; 25 °C; 60 min; 40 kHz
2.5 mg g–1
Hydroxyl DPPH
Stojičević et al., 2008
LSR 1:10; 5%NaOH; 60 °C; 60 min; 100 W
24.3% polysaccharides
DPPH
Hromádková et al., 2008
Sappan wood (Caesalpinia sappan)
540 W; 20 min
0.747 g
DPPH Nitric oxide
Shrishailappa et al., 2007
Cranberry (Vaccinium macrocarpon)
LSR 1:5.7; 10 min; 125 °C W E A LSR 1:50; 70% M; 50 W; 10 min
Ginseng (Panax ginseng) Ginseng (Panax ginseng) Indian mulberry (Morinda citrifolia)
LSR 1:50; 70% E; 250 W; 10 min 2450 MHz LSR 1:100; E; 720 W; 30 min; 60 °C
17.6% (9.42 mg Q g–1) 2% (1272 mg Q g–1) 3.25% (960 mg Q g–1) G: 9.8 mg TAE g–1 dw TP: 5.8 mg TAE g–1 dw TS: 3.30% 65.88%
Inhibition of lipid Raghavan and oxidation Richards, 2007 TBARS DPPH
Kim et al., 2007
DPPH
Chen et al., 2010
DPPH
Hemwimol et al., 2006
554 Separation, extraction and concentration processes
Table 18.11 Continued
141 W; 83 s; 9.8 mL; 4 cycles
5.94 mg A mg–1
DPPH
Zhao et al., 2009
LSR 1:10; E 95%; 30 min; 500 W; 2450 MHz, 80 °C
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ACN, acetonitrile; E, ethanol; LSR, liquid:solid ratio; M, methanol; RT, room temperature; W, water; dW, distilled water. A, astaxanthin; CE, catechin equivalent; FG, flavanones glycosides; G, ginsenosides; GAE, gallic acid equivalent; Q, quercetin; TAE, tannic acid equivalent; TCE, total content of extract; TP, total phenolics; TEP, total extractable phenolics; TS, total saponins.
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Green algae (Haematococcus pluvialis) Longan peel (Dimocarpus longan Lour.) Teja (Cinnamomum iners Reinw.) Rice bran (Oryza sativa)
556 Separation, extraction and concentration processes et al., 2008). PEF enhances mass transfer by improving softness and texture (Eshtiaghi and Knorr, 2002), and increasing the contents of some solutes, including total phenolics. However, the antioxidant activity varies significantly with the treatment used (Corrales et al., 2008). PEF inactivates degrading enzymes. HHP increases the extraction yields by deprotonation of charged groups, disruption of salt bridges and hydrophobic bonds in cell membranes and decreasing the dielectric constant of water (Corrales et al., 2008). Combinations of the different extraction methods are used. High-pressure extraction combined with sonication was used to extract compounds from Berberis koreana bark, improving the extraction of bioactive compounds compared with the high-pressure process (Qadir et al., 2009).
18.6 Integration of extraction processes and purification Several technologies with different degrees of sophistication (including solvents, resins and membranes) were used to obtain purified compounds or fractions with biological activity from crude extracts. Some examples of their individual and combined application to the concentration and purification of natural antioxidants from several sources are included here. Combinations of conventional and emerging extraction technologies were proposed to extract and/or to purify natural compounds with antioxidant and biological activity. The extraction of oil and/or lipophylic fractions with supercritical fluids at low pressures and the subsequent solvent extraction of the solid residue was proposed to recover essential oils from aniseed (Rodrigues et al., 2003), Melissa officinalis (Marongiu et al., 2004), rosemary, sage (Nakatsu and Yamasaki, 2000), summer savory (Esquível et al., 1999), and oils from borage seeds (Soto et al., 2008). A sequence of SFE and hydrothermal extraction has been proposed to obtain antioxidants from bamboo (Quitain et al., 2004). Alternatively, a first extraction of the raw material with conventional solvents and further purification of the crude extracts by SFE offer operational advantages, including reduction in CO2 consumption, increased extraction yields, and processing of less substrate owing to its enhanced content of bioactive compounds. Two options are used to purify the bioactive compounds from the crude extract: relatively mild SFE to extract non-polar compounds (desirable when antioxidant compounds of a phenolic nature have to be recovered from the residue) (Hadolin et al., 2004), and SFE under harsh conditions to recover polar compounds in the extract (Díaz-Reinoso et al., 2006). A simple purification process based on the sequential extraction of the raw materials or extracts with solvents of different polarity was applied to the extraction of antioxidant compounds from plant foods. Cheung et al. (2003) reported on the processing of edible mushrooms, leading to similar yields as those obtained with methanol and water on large and small scales. The
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Extraction of natural antioxidants from plant foods 557 chemical components of the methanol extract were fractionated by solvents of increasing polarity (dichloromethane, ethyl acetate, butanol and water) (Cheung and Cheung, 2005). The dichloromethane subfraction of the methanol extract of Volvariella volvacea had a high antioxidant activity against lipid peroxidation of rat brain homogenate, whereas the ethyl acetate subfraction of the methanol extract of V. volvacea has antioxidant activity against the oxidation of human low density lipoprotein (LDL) similar to that of caffeic acid. Combined extraction involving pressurized hot water extraction as a first stage has been reported for a variety of purposes (Chambers et al., 1984; Wang and Weller, 2006). The advantages of membrane processing include a low energy requirement, no additives, mild operating conditions, separation efficiency and easy scaling up. The utilization of membrane technologies for concentrating and purifying bioactive phenolic compounds from aqueous streams is a topic of growing interest. For example, fractionation of phenolics into low- and high- MW compounds was reported for conventional solvent crude extracts from edible mushrooms (Cheung and Cheung, 2005), grape (Nawaz et al., 2006), grape pomace (Díaz-Reinoso et al., 2009), partially purified extracts from persimmon pulp (Ga et al., 2008), and mulberry root cortices (Yu et al., 2007). Ultrafiltration (UF) membranes are easy and fast to use in the separation of phenolics according to their MW. At the analytical scale, almond skin phenolics were fractionated to separate low MW compounds in permeate and proanthocyanidin oligomers (up to decamers) in retentate. Permeates from a 50 kDa membrane contained proanthocyanidin pentamers (Prodanov et al., 2008). UF membranes can be used for tailoring grape anthocyanins (containing about 60% monomers, 20% polymers, and 20% other forms) from grapes, having benzoic acids, hydroxycinnamates, anthocyanins, flavan-3-ols, and flavonols as major phenolic groups. The membrane (<100 kDa) separated the polymeric form in the retentates, wheareas permeates contained low-MW compounds. After a 10-fold reduction in volume, the permeate had higher monomer and lower polymer content than the initial feed, with the opposite for the retentate. The composition of the fractions affected color, functional properties and antioxidant activity. The total phenolic content correlated linearly with the antioxidant activity, whereas lightness and color properties related linearly to the monomeric content (Kalbasi and Cisneros-Zevallos, 2007). Ultrafiltration (0.22 mm) of 50% ethanol solvent of grape seeds was employed to maximize the recovery of polyphenols (Nawaz et al., 2006), whereas nano- and ultrafiltration were used to process aqueous extracts from grape pomace (Díaz-Reinoso et al., 2009). Membrane separation was applied to the extracts from both conventional and alternative extraction processes. Samples of edible mushrooms (Lentilula edodes, V. volvacea) were extracted sequentially with petroleum ether, ethyl acetate and methanol, and the insoluble residue was first extracted with
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558 Separation, extraction and concentration processes boiling water and further processed by UF (10 kDa) to yield fractions of low and high MW. The low-MW fraction was highly active against lipid peroxidation of rat brain homogenate (Cheung and Cheung, 2005), and the activity seemed to correlate with their content of protein/free amino acids. SFE, ultrasound-aided extraction and membranes were applied to the manufacture of algal polysaccharides. Sargassum pallidum powder was extracted by SFE and the resulting powder was disrupted by ultrasonic waves, incubated, and centrifuged. Proteins in the supernatant solution were removed by adding trichloroacetic acid, and the resulting solution was adjusted to pH 7, passed through membranes having different MW cut-offs, and subjected to rotary evaporation, ethanol precipitation and freeze-drying. Crude polysaccharides obtained from UF were dissolved and applied to a column of DEAE-52 cellulose to obtain seven fractions of polysaccharides (Ye et al., 2008). Among them, the low-MW fractions with high sulfate content presented the highest antitumor activity against HepG2 cells, A549 cells, and MGC-803 cells (Ye et al., 2008). SFE was used for purification of tea polyphenols as part of a complex sequence involving ultra- and nanofiltration (Fig. 18.3), vacuum concentration, spray drying and conventional solvent extraction. A 1000-Da membrane was used to process spinach extract produced by pressurized liquid water or by ethanol extraction, yielding a high-MW fraction with much higher antioxidant activity than that of the low-MW fraction (Howard and Pandjaitan, 2008). The separation selectivity is improved by using membranes able to interact with some compounds in solution. For example, a modified PVP membrane favoring the formation of hydrogen bonds with flavonoids improved their separation from G. biloba extracts (Xu et al., 2005b). UF was used in concentrating melanoidins from coffee brew. Operating in discontinuous diafiltration mode (Rufián-Henares and Morales, 2007), the retentates contained soluble melanoidins with antioxidant activity, whereas permeates contained compounds non-covalently bound to melanoidins, which showed no antioxidant activity (Delgado-Andrade and Morales, 2005). Novel technologies (such as membrane distillation and osmotic evaporation) were proposed to produce juice concentrates preserving their natural quality and antioxidant activity (Cassano et al., 2007; Koroknai et al., 2008) (Fig. 18.4). The combination of protease hydrolysis and membrane fractionation of peptides is an attractive method to improve the potential of proteins with high nutritive value but poor functional properties (e.g. solubility, emulsification and film-forming properties) and negative sensory properties (color, taste and texture). The functional and immunological properties of proteins can be modified, enabling their utilization in applications such as additives for beverages, hypoallergenic infant diets, nutritional therapies, food texture enhancers, or pharmaceutical ingredients. The functionality of the final product can be controlled by selecting specific enzymes and reaction factors (Cheryan
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Fig. 18.3 Processing scheme involving supercritical CO2, and ultra- and nanofiltration.
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Extraction of natural antioxidants from plant foods 561 and Deeslie, 1983; Deeslie and Cheryan, 1988). Fractionation of protein hydrolyzates is a typical area for membrane applications. Examples dealing with conventional and non-conventional protein sources are presented. Rapeseed protein isolates are suitable for food applications owing to their functional properties and health benefits, such as ACE inhibition, bile acidbinding and free radical-scavenging activity. Since the antinutritive and toxic compounds (such as glucosinolates and phytates) have significantly lower MW than rapeseed proteins, they can be selectively separated by precipitation at controlled pH or removed by UF. The ultrafiltered protein extract had good functional properties, whereas the precipitated protein showed stronger ACE inhibition, bile acid-binding capacity and DPPH radical-scavenging capacity (Yoshie-Stark et al., 2008). Soybean peptides, with or without enzymatic treatment, were selectively retained by UF membranes (Moure et al., 2006) (Fig. 18.5). Alfalfa leaf protein was hydrolyzed with protease, and the reaction products, fractionated by UF and purified by adsorption, showed high nutritive value, chelating ability, reducing power, and radical scavenging activity. Mouse breeding with these products resulted in increased activities of glutathione peroxidase and superoxide dismutase, and decreased the formation of malonaldehyde by oxidative reaction (Xie et al., 2008). Proteins extracted from potato tubers and by-products from the potato industry were hydrolyzed and ultrafiltered to yield a permeate containing ACE-inhibitory compounds (Pihlanto et al., 2008). Papain was used for preparing hydrolyzates from wheat gluten, which were separated by UF to yield fractions with strong antioxidative activities measured by the linoleic acid and DPPH tests (Wang et al., 2007). The hydrolyzates produced by enzymatic treatment of wheat gluten with commercial proteases were fractionated by UF, to yield products with properties strongly related to their molecular weight distribution and amino acid composition. The activities of selected fractions were measured by the linoleic acid and scavenging radical assays (Kong et al., 2008). UF membrane reactors have several advantages over conventional technologies, including enzyme reuse and accurate selection of the product molecular size. The flowsheet of a process for producing soy protein hydrolyzate using a continuous UF membrane reactor fed with a combination of commercial enzymes is shown in Fig. 18.6. The amino acid sequences of peptides depended on the specificity of enzymes, whereas the characteristics of the membrane governed the functional properties of protein hydrolyzates (Chiang et al., 1999). The major polysaccharides from olive fruits are pectins, which are esterlinked to phenolic acids. The action of selected enzymes facilitates the release of phenols. Bouzid et al. (2005) proposed the incubation with coumaroyl ester hydrolase and further processing to recover hydroxytyrosol and free phenolic acids (caffeic, p-coumaric and ferulic acids). Adsorption onto polymeric resins was successfully used for the separation
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and purification of phenolic compounds from natural products (Kammerer et al., 2005; Llorach et al., 2004; Saleh et al., 2008; Scordino et al., 2005). Separation and purification of rutin from leaves, flowers, stalks, and seeds of buckwheat (Fagopyrum esculentum Moench) has been reached by a lowcost process consisting of extraction with aqueous 50% ethanol at 80 °C for 1 h, adsorption onto a styrene-based resin (HP-20), elution with water and aqueous ethanol (20% and 30%), and recrystallization at 4 °C for 12 h). Rutin recovery was 92%, with over 95% purity (Kyoung et al., 2005). Extraction of dried apple pomace (a by-product of apple juice production) with diluted mineral acid, followed by adsorption of phenolics (phloridzin, chlorogenic acid and quercetin glycosides) was carried out using a hydrophobic styrene–divinylbenzene resin (Amberlite XAD 16HP). Pectins were eluted with distilled water, and the phenolic compounds with methanol (Schieber et al., 2003). Water and alcoholic extracts from lettuce, chicory by-products and cauliflower (Brassica oleracea L. var. botrytis) were purified by adsorption on Amberlite XAD-2. The antioxidant capacity (reducing power and DPPH, ABTS radical scavenging capacity) and the capacity to inhibit lipid peroxidation (ferric thiocyanate assay) were linearly correlated with the phenolic content (Llorach et al., 2003 and 2004). Polysaccharide fractions prepared from various materials are usually prepared by sequential steps of solubilization, precipitation and fractionation or purification. For example, crude extracts from Cuscuta chinensis seeds were extracted with hot water and diluted alkali, and then precipitated by addition of ethanol. Anion exchange and gel filtration chromatography were used for purification of an acidic polysaccharide with rhamnogalacturonanlike structure (Bao et al., 2002). Alkaline extraction of ferulic acid from © Woodhead Publishing Limited, 2010
564 Separation, extraction and concentration processes maize bran was carried out using 2M NaOH. Ferulic acid was purified by adsorption chromatography followed by preparative high-performance thinlayer chromatography, yielding a product with a purity of 95.35% (Tilay et al., 2008). Novel adsorbents for polyphenols include silk fibroin, which has been used to process ethanol extracts from olive leaves. This material adsorbed 15 mg g–1 rutin and 96 mg g–1 oleuropein, as well as other polyphenols such as verbascoside, apigenin-7-glucoside, and luteolin-7-glucoside (Altiok et al., 2008). Woody materials such as agricultural by-products or natural fibres are useful for isolating bioactive components from plants. Substrates prepared from woody tea stalk, pine sawdust and sugarcane bagasse were suitable to adsorb decaffeinated catechins from tea extracts. The tea components were separated by gradient elution with increasing ethanol concentration, yielding fractions with different proportions of caffeine, partial non-gallated catechins and gallated catechins. The concentration of total catechins in the selected ethanol eluates was above 90%, with caffeine below 1% (dry basis) (Ye et al., 2009). The water-soluble fraction resulting from steam treatment of olives (Fernández-Bolaños et al., 2002) contains other high-added-value compounds such as monosaccharides, oligosaccharides, and mannitol. Oligosaccharides were separated by size-exclusion chromatography, whereas highly purified mannitol could be recovered by a simple purification method (FernándezBolaños et al., 2004). Highly purified 3,4-dihydroxyphenyl glycol (DHPG) was obtained from alperujo by chromatography to yield 96% pure product at 21% yield. From 1000 kg of wet alperujo (300 kg of dry matter), 807 g DHPG could be obtained (Rodríguez et al., 2009). Hydroxytyrosol was recovered from olive cake at good yield (1–1.2 g hydroxytyrosol/100 g of dry matter) upon acidic processing (Fernández-Bolaños et al., 2002). Crude rapeseed peptides and peptide fractions with antioxidant and antithrombotic activities were prepared by incubation with enzymes and further adsorption. The rapeseed slurry from a wet-milling was treated with a combination of enzymes, and subjected to thermal inactivation and centrifugation. The aqueous phase was adjusted to pH 4 with acetic acid, adsorbed in the macroporous adsorption resin and eluted with deionized water and ethanol. Stepwise desorption with 25–55% ethanol enabled the recovery of different fractions. The fraction eluted with 55% ethanol showed more potent antioxidant activities (reducing power, inhibition of lipid oxidation in a liposome model) except for hydroxyl radicals, probably owing to the higher contents of hydrophobic amino acid, tannin, and the brown color substances. A positive correlation existed between the peptide concentration and antioxidant activity (Zhang et al., 2007). Tricin (5,7,4¢-trihydroxy-3¢,5¢-dimethoxyflavone), a cancer chemopreventive agent, was prepared from a concentrated extract from bamboo leaves and further processed by adsorption in a polystyrene resin, preparative high-
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Extraction of natural antioxidants from plant foods 565 performance liquid chromatography, dialysis and crystallization (Jiao et al., 2007), resulting in higher yields and fewer processing steps than the normal synthetic approach. Membranes and resins were proposed for processing hydrolyzates from alkaline treatments. Liquors from the alkaline processing of sugarcane bagasse were ultrafiltered, and the phenolic acids present in permeates were bound to an anion-exchange resin. After washing, elution with a mixture of water–ethanol–HCl followed by crystalization and washing of crystals with 1% HCl led to a concentrate containing products with 89.7% purity (measured as coumaric acid), which had the same antioxidant activity, reducing power and free radical scavenging capacity as the standard p-coumaric acid (Ou et al., 2009). Combinations of different treatments were also applied to the extraction and purification of polysaccharides from Physalis alkekengi var. francheti fruit, including hot water extraction, ultrasonic-assisted extraction and enzyme pretreatment. Fractionation of polysaccharides with DEAE and Sephadex G-200 led to fractions with different radical scavenging activity (Ge et al., 2009). Membranes and resins have been combined in processing schemes similar to that given in Fig. 18.7. A macroporous adsorption resin was used to remove low-MW polar substances (including sugars, gallic acid, and organic acids) and to increase the contents of total polyphenols and condensed tannins. The partially purified extract was subjected to UF (10 kDa) to yield a retentate containing condensed tannins of high MW, whereas permeate contained lowMW phenolic compounds. High-MW tannins showed high hydroxyl radical scavenging activity, as a result of their structure. This fraction also exhibited dose-dependent, inhibitory activity against hydroxyl and superoxide anion radical scavenging activity, as well as activity in the peroxidation of linoleic acid. Low-MW tannins were reported to be pro-oxidant during accelerated lipid peroxidation (Gu et al., 2008). UF of green tea water extracts through a composite membrane was suitable for retaining particles, proteins, polysaccharides and tannic acid, leading to a permeate containing more than 40% phenols (epicatechin, epicatechin gallate, epigallocatechin, epigallocatechin gallate) and caffeine (Li et al., 2005). This stream was further separated using adsorption–desorption resins, to yield a purified product containing more than 90% polyphenols (Li et al., 2005). A combination of membrane fractionation and resin purification was also applied to peptide purification. A recent example is the purification of nucleoprotein complexes isolated from Saccharomyces cerevisiae by mild alkaline extraction and precipitation with acetic acid. The high- and low-MW fractions of the nucleoprotein were separated by cross-flow microfiltration (Butylina et al., 2007). S. cerevisiae cells were exposed to hydrogen peroxide or to ultraviolet irradiation to induce oxidative∑ stress: the first acts as an oxidant by producing hydroxyl free radicals (OH ) that attack cell components and cause various modifications in human cells, whereas the latter can cause
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Extraction of natural antioxidants from plant foods 567 damage to nucleic acids and proteins, and affect other molecules through production of ROS. The addition of a high-MW nucleoprotein complex to damaged cells enhanced growth rate of the yeast population up to a similar value to that of intact cells (Butylina et al., 2007).
18.7 Future trends Antioxidants are commonly used in foods to retard or prevent deterioration via lipid oxidation, which leads to the development of undesirable rancid odors, off-flavors, discoloration and generation of potentially toxic compounds, limiting the quality, acceptability and shelf life of processed products. As a general trend, obtaining a similar degree of protection as synthetic antioxidants requires higher doses of natural extracts. However, if the extracts lack toxicity and add functional and biological properties to the product, higher loadings could be used. The use of natural antioxidants would be better restricted to those cases in which their application is necessary, and the use of materials relatively rich in antioxidants should be preferred to the application of extracts, concentrates or pure mixtures of active components. In this context, the utilization of food processing wastes as feedstocks could facilitate the production of valuable natural products, which would guarantee both sustainability and satisfaction of consumer demands. In the future, the application of advanced, efficient technologies such as ultrasound, pressurized extraction, and SFE is expected to offer great potential and selectivity for process development. Technological advances in these methods would require specific optimization. Compared with conventional extraction methods, the combination of low cost raw materials and effective extraction technologies is of environmental and economical interest. These techniques do not involve the utilization of organic solvents, allow reduced extraction times and may improve separation selectivity.
18.8 Sources of further information and advice Owing to the wide-ranging nature of this work, the set of references included provide a non-exhaustive overview of the field. Because of this, special attention has been devoted to include recent reviews, dealing with specific topics, which can be of help for futher information. However, the intensive research on antioxidants is reflected in a high rate of publication, particularly in topics related to the bioactivity of selected compounds suitable for key applications such as dermatological preparations, cancer, cardiovascular and neurodegenerative diseases.
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18.9 Acknowledgements The authors are grateful to the Spanish Ministry of Education and Science (MEC) for the financial support of this work (in the framework of the Research Project reference ALG2006-05387, which had partial financial support from the FEDER funds of the European Union).
18.10 References Adil I H, Yener M E and Bayindirli A (2008), ‘Extraction of total phenolics of sour cherry pomace by high pressure solvent and subcritical fluid and determination of the antioxidant activities of the extracts’, Separation Science and Technology, 43, 1091–1110. Aehle E, Raynaud-Le Grandic S, Ralainirina R, Baltora-Rosset S, Mesnard F, Prouillet C, Mazière J-C and Fliniaux M-A (2004), ‘Development and evaluation of an enriched natural antioxidant preparation obtained from aqueous spinach (Spinacia oleracea) extracts by an adsorption procedure’, Food Chemistry, 86, 579–585. Ahn G N, Park E J, Kim D S, Jeon Y J, Shin T K, Park J W, Woo H C, Lee K W and Jee Y (2008), ‘Anti-inflammatory effects of enzymatic extract from Ecklonia cava on TPA-induced ear skin edema’, Food Science and Biotechnology, 17, 745–750. Ahn M-J, Yoon K-D, Min S-Y, Lee J S, Kim J H, Kim T G, Kim S H, Kim N-G, Huh H and Kim J (2004), ‘Inhibition of HIV-1 Reverse transcriptase and protease by phlorotannins from the brown alga Ecklonia cava’, Biological & Pharmaceutical Bulletin, 27, 544–547. Almeida I F, Valentão P, Andrade P B, Seabra R M, Pereira T M, Amaral M H, Costa P C and Bahia M F (2008), ‘In vivo skin irritation potential of a Castanea sativa (chestnut) leaf extract, a putative natural antioxidant for topical application’, Basic & Clinical Pharmacology & Toxicology, 103, 461–467. Alonso-Salces R M, Korta E, Barranco A, Berrueta L A, Gallo B and Vicente F (2001), ‘Determination of polyphenolic profiles of Basque cider apple varieties using accelerated solvent extraction’, Journal of Agricultural and Food Chemistry, 49, 3761–3767. Altiok E, Baycin D, Bayraktar O and Ülkü S (2008), ‘Isolation of polyphenols from the extracts of olive leaves (Olea europaea L.) by adsorption on silk fibroin’, Separation and Purification Technology, 62, 342–348. Altunkaya A, Becker E M, Gökmen V and Skibsted L H (2008), ‘Antioxidant activity of lettuce extract (Lactuca sativa) and synergism with added phenolic antioxidants’, Food Chemistry, 115, 163–168. Amaral S, Mira L, Nogueira J M F, da Silva A P and Florencio M H (2009), ‘Plant extracts with anti-inflammatory properties – a new approach for characterization of their bioactive compounds and establishment of structure – antioxidant activity relationships’, Bioorganic & Medicinal Chemistry, 17, 1876–1883. Amarowicz R and Shahidi F (1997), ‘Antioxidant activity of peptide fraction of capelin protein hydrolysates’, Food Chemistry, 58, 355–359. Ames B N, Shigenaga M K and Hagen T M (1993), ‘Oxidants, antioxidants, and the degenerative diseases of aging’, Proceedings of the National Academy of Sciences of the United States of America, 90, 7915–7922. Andrade P B, Pereira D M, Ferreres F and Valentao P (2008), ‘Recent trends in high throughput analysis and antioxidant potential screening for phenolics’, Current Pharmaceutical Analysis, 4, 137–150. Andreasen M F, Christensen L P, Meyer A S and Hansen A (1999), ‘Release of hydroxycinnamic and hydroxybenzoic acids in rye by commercial plant cell wall
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Fractionation of egg proteins and peptides 595
19 Fractionation of egg proteins and peptides for nutraceutical applications B. P. Chay Pak Ting, Y. Pouliot and S. F. Gauthier, Laval University, Canada and Y. Mine, University of Guelph, Canada
Abstract: The main components of egg proteins and their physicochemical characteristics, and bioactive peptides derived by enzymatic hydrolysis from egg proteins, are described. An overview of recent developments in fractionation and purification processes for bioactive proteins/peptides from egg proteins is also presented. Further developments of techniques required to achieve the separation of specific proteins/peptides are described and the potential for creating new value-added ingredients with applications in the food, nutraceutical and biotechnological industry is explored. Key words: egg proteins, proteins, peptides, separation, bioactivity.
19.1 Introduction The avian egg is considered to be a rich source of nutrients, such as proteins, lipids, and enzymes, and biological substances, such as growth promoting factors and defence factors. Overall, an egg is constituted of 63% egg white, 27.5% egg yolk and 9.5% eggshell (Table 19.1). Although eggs contain about 75% water, they are a rich source of high-quality protein, unsaturated fatty acids, vitamins and minerals. Egg proteins are essentially distributed between the egg white and the yolk, with a small proportion in the eggshell. Lipids, negligible in egg white, are almost exclusively found in the egg yolk and are associated with proteins to form lipoproteins. Carbohydrates are minor egg components, which are present in the egg as both free carbohydrates and bound to proteins/lipids. Most of the minerals are found in the eggshell and in the yolk, where phosphorus and potassium are in soluble form.
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596 Separation, extraction and concentration processes Table 19.1 Chemical composition of hen eggs and their major components [data from Li-Chan and Nakai (1989)] Constituent
Egg shell Egg white Egg yolk
% (w/v)
9.5 63.0 27.5
Major components (%, w/w) Protein
Lipid
Carbohydrate
Ash
6.4 9.7–10.6 15.7–16.6
0.03 0.03 31.8–35.5
– 0.4–0.9 0.2–1.0
– 0.5–0.6 1.1
Table 19.2 Composition of egg white and yolk proteins
Egg white
% of total proteins
Ovalbumin Ovotransferrin Ovomucoid Ovomucin Lysozyme G2 globulin G3 globulin Ovoinhibitor Ovoglycoprotein Ovoflavoprotein Ovomacroglobulin Cystatin Avidin
54 12 11 3.5 3.4 4.0 4.0 1.5 1.0 0.8 0.5 0.05 0.05
Lipovitellin Phosvitin LDLg Livetin LDL
70 16 12 15 85
Egg yolk Granules Plasma
LDLg, low-density lipoprotein from granules.
Table 19.2 lists the major proteins of egg white and egg yolk together with their average concentration. The major proteins in egg white are ovalbumin, ovotransferrin, ovomucoid, ovomucin and lysozyme which account for >83% of total egg white proteins. The major egg yolk proteins take the form of lipoprotein complexes, which are divided into the plasma and granule fractions comprising lipovitellin, phosvitin, livetin and lipovitellenin. A number of egg proteins and some of their peptide sequences that have biological activities have been identified. For example, lysozyme is the egg protein that has probably attracted the most attention because of its potential use as an antimicrobial agent in foods. Chicken lysozyme contains a peptide sequence that is potently antimicrobial against both Gram-positive and Gram-negative bacteria. Biologically active proteins and peptides occur at relatively low concentrations in foods and pre-concentration of purification © Woodhead Publishing Limited, 2010
Fractionation of egg proteins and peptides 597 is needed to obtain a dose level to produce beneficial effects in situ. It was therefore necessary to develop technologies for fractionation and purification of bioactive molecules of interest for nutraceutical applications. This chapter mainly focuses on the chemistry and biological activities of proteins and peptides derived from egg components and the various technologies applied for fractionation/purification of proteins/peptides.
19.2 Composition and physicochemical characteristics of egg proteins and applications in the nutraceutical industry 19.2.1 Egg white The physicochemical characteristics of some egg white proteins are listed in Table 19.3. Egg white proteins are predominantly globular proteins having an acidic isoelectric point (pI), the exceptions being lysozyme and avidin. Except for lysozyme, egg white proteins are glycoproteins that are rich in sulfur amino acids and very sensitive to heat denaturation. Ovalbumin Ovalbumin has a molecular mass of 45 kDa and is composed of 385 amino acids. This protein belongs to the serpins family and possesses two genetic variants which differ at residue 289 and 311 by a substitution of Glu and Asn, respectively, by Gln and Asp (Ishihara et al., 1981; Wiseman et al., 1972). The primary sequence reveals one disulfide bond between Cys74 and Cys121. Table 19.3 Physicochemical characteristics of egg white proteins [adapted from LiChan and Nakai (1989)]
Amino acid residues
MW (kDa)
pI
T d*
Number of cystein residues [disulfide bridge]
Ovalbumin Ovotransferrin Ovomucoid Ovomucin Lysozyme G2 globulin G3 globulin Ovoinhibitor Ovoglycoprotein Ovoflavoprotein Ovomacroglobulin Avidin Cystatin
385 686 186 8300 129 – – – – 219 – 128 –
45 77.7 28 5500–8300 14.3 30–45 – 49 24.4 32 760–900 68.3 12.7
4.5 6.0 4.1 4.5–5.0 10.7 5.5 4.8 5.1 3.9 4.0 4.5 10 5.1
84 61 79 – 75 92.5 – – – – – 85 –
6 [1] 30 [15] 18 [9] – 8 [4] – – – – 18 [9] – 2 [1] 4 [2]
(–) Data not available. * Denaturation temperature.
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598 Separation, extraction and concentration processes Approximately one half of ovalbumin amino acids residues are hydrophobic and one third are acidic, which confer a pI of 4.5 (Li-Chan and Nakai, 1989). The secondary structure of ovalbumin consists of 30% a-helix and 32% b-sheet structure using high resolution x-ray diffraction (Stein et al., 1990, 1991). Three components A1, A2 and A3 have two, one and no phosphate, respectively, and are found in purified ovalbumin (Perlman, 1952). In solution, ovalbumin can be denatured and aggregated by thermal denaturation or by exposure to the air–water surface (Mine et al., 1990). Partial thermal denaturation of ovalbumin occurs at 78 to 86 °C. The protein adopts a more stable form, termed S-ovalbumin to denote its increased stability. At a heating rate of 10 °C min–1 at pH 9, the denaturation temperature of ovalbumin is 84.5 °C compared with 92.5 °C for S-ovalbumin (Donovan and Mapes, 1976). The greater stability, compactness and hydrophobicity of the S-form contrast with that of ovalbumin (Nakamura and Ishimaru, 1981). Ovotransferrin Ovotransferrin, also called conalbumin, belongs to the transferrin family. Ovotransferrin, with a molecular mass of 77.7 kDa and 686 amino acids, contains no phosphorus or free sulfhydryl group. Ovotransferrin has a pI of 6.0 and has the capacity to bind two Fe3+ ions per molecule with two CO2–3 or HCO3– ions (Mine, 2007). Ovotransferrin is folded into two lobes which show similar structural elements (Williams et al., 1982). Each lobe contains a site for iron binding and is further divided into two distinct domains. Ovotransferrin is the most heat-sensitive protein of egg white. The protein easily binds metallic ions, such as Fe3+, Al3+, Cu2+ or Zn2+ forming heatstable complexes. Ovomucoid Ovomucoid is a glycoprotein with a molecular weight of 28 kDa and a pI of 4.1. Its amino acid sequence is composed of 186 residues and it possesses no tryptophan residues. Ovomucoid has nine disulfide bridges and no sulfhydryl groups. Ovomucoid has three domains defined by the amino acid sequences of 1–68, 69–130 and 131–186, each domain is cross linked by three disulfide bridges (Kato et al., 1987). Its secondary structure is composed of 26% a-helix, 46% b-structure, 10% b-turn and 18% random coil structure (Watanabe et al., 1981). Ovomucoid is highly resistant to heat owing to its high cystine content and, consequently, to the high number of disulfide linkages. Under acidic conditions ovomucoid can resist to heat treatments up to 100 °C, but it is rapidly (89 °C) denaturated in alkaline solutions (pH 9) (Matsuda et al., 1982). Ovomucin Ovomucin is a highly glycosylated protein displaying a molecular weight in the range of 5500–8300 kDa depending on the method of isolation and environmental conditions (Li-Chan and Nakai, 1989). Ovomucin differs from
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Fractionation of egg proteins and peptides 599 other egg white proteins because it is extremely large. Its structure comprises sulfate esters, large amounts of cystine and 50% of the total sialic acid contents of egg white (Stadelman and Coterril, 1973). Ovomucin is composed of two subunits: a- and b-ovomucin, which differ by a carbohydrate-poor fraction (15% of dry matter) and a carbohydrate-rich fraction (50% of dry matter), respectively, and amino acids composition. Watanabe et al. (2004) estimated the molecular weight of a-ovomucin at 254 kDa whereas that of b-ovomucin has been estimated, using SDS-PAGE, to be between 400 and 720 kDa (Itoh et al., 1987). Ovomucin is heat stable over the pH range between 7.1 and 9.4. Cunningham and Lineweaver (1965) showed that the viscosity and absorbance of ovomucin solutions did not change upon heat treatment at 90 °C for 2 h. Lysozyme Lysozyme is a small enzyme molecule of 14.3 kDa capable of hydrolyzing specific polysaccharides that cleaves the b 1-4 linkages between N-acetylneuraminic and N-acetylglucosamine in bacteria cell. The single polypeptide chain consisting of 129 amino acid residues is cross-linked by four disulfide bridges. Because of its strong basic character (pI = 10.7), lysozyme binds to ovomucin, ovotransferrin and ovalbumin, via electrostatic interactions between positively charged lysozyme and negatively charged residues of sialic acid in these glycoproteins. Young et al. (1994) analyzed the three-dimensional structure of hen egg white lysozyme. The lysozyme molecule has two domains: the N-terminal domain (residues 40 to 88) with a hydrophobic core is composed of antiparallel b-sheets whereas the second hydrophilic domain, residues 1 to 39 and 89 to 129, is made up of a-helix. The enzyme is much more heat sensitive in egg albumen than when present alone. Minor egg white proteins The remaining proteins listed in Table 19.2 and 19.3 (G2 and G3 globulin, ovoinhibitor, ovoglycoprotein, ovoflavoprotein, ovomacroglobulin, cystatin, avidin) account for <17% of egg white proteins. Although some of these proteins may have potential applications as bioactive proteins or peptides, their sequence and structure and properties have not been fully characterized to date. 19.2.2 Egg yolk Egg yolk is a complex mixture of micro granules held in suspension. Proteins and lipids are the main constituents of yolk accounting for 15.7–16.6 and 31.8–35.5%, respectively. The yolk lipid fraction contains 66% triacylglycerol, 28% phospholipid, 5% cholesterol and minor amounts of other lipids (Powrie and Nakai, 1985). Egg yolk can be separated into plasma (the supernatant) and granule (the precipitate) after dilution in a saline solution (0.17 M
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600 Separation, extraction and concentration processes NaCl) followed by a centrifugation. Plasma is composed of 85% low-density lipoproteins (LDL) and 15% livetin and granules contain mainly 70% lipovitellin (HDLs), 16% phosvitin and 12% LDLg (Burley and Cook, 1961). HDLs and phosvitin come from the same precursor, namely vitellogenin II, and they are linked by phosphocalcic bridges which probably form the basic unit of the granules (Radomski and Cook, 1964). After uptake by the oocyte, vitellogenin undergoes specific enzymatic cleavage to generate lipovitellin I (120 kDa) and lipovitellin II (30 kDa) in the N-terminal and C-terminal region, respectively, with the phosphoseryl-rich domain termed phosvitin lying in between. Protein composition of egg yolk granules Phosvitin Egg yolk phosvitin is a phosphoglycoprotein with a molecular weight of 35 kDa containing 10% phosphorus and 6.5% carbohydrates (Mecham and Olcott, 1949). Phosvitin is constituted of 217 amino acid residues which comprise a core region of 99 amino acids, consisting of 80 serines, grouped in runs of maximally 14 residues interspersed by arginines, lysines and asparagines. Most of serine residues are phosphorylated and the phosphoserines are forming blocks that can carry up to 15 consecutive residues (Byrne et al., 1984; Van Het Schip et al., 1987). The relative abundance of phosphoseryl groups in the phosvitin amino acid sequence confers to the protein a large central hydrophilic portion surrounded by two small hydrophobic parts at the N-terminal and C-terminal. Owing to its polyanionic character (pI = 4), phosvitin possesses a very strong metal-chelating property. It can bind multivalent metals and 95% of Fe in egg yolk is complexed together with phosvitin (Greengard et al., 1964). Castellani et al. (2004) have found that pH 6.5 and ionic strength of 0.15 M were optimal for iron binding by phosvitin. Fourier transform infrared spectroscopy showed that the secondary structure of phosvitin is composed of 0% a-helix, 50% b-sheets, 7% b-turns and 43% random coil (Losso et al., 1993). Two constituents from hen’s egg yolk phosvitin, namely a-phosvitin and b-phosvitin with molecular weight of 160 000 and 190 000 Da, respectively, migrate in glycin-2,6-lutidine buffer without sodium dodecylsulfate (SDS) using electrophoresis. Under denaturating conditions with 0.5% SDS and Tris–glycine buffer, a-phosvitin and b-phosvitin dissociate into polypeptides with a molecular weight of 37.5, 42.5, 45 and 45 kDa, respectively (Abe et al., 1982). They also differ by their amino acid compositions, phosphorus contents, concentrations in carbohydrates, and solubility in presence of CaCl2 (Itoh et al., 1983). Phosvitin is relatively heat stable. Precipitation of phosvitin solution does not occur after heating at 100 °C between pH 4 to 7 (Mecham and Olcott, 1949). Itoh et al. (1983) observed no change in the electropherogram of aand b-phosvitin heated at 110 °C but the phosvitin bands were completely diffused at 140 °C.
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Fractionation of egg proteins and peptides 601 Lipovitellins (HDLs) In native egg yolks, lipovitellins or HDLs are linked to phosvitin to form the granular unit through phosphocalcic bridges. HDLs are in form of dimer of 400 kDa and are made up of 75–80% proteins and 20–25% lipids. HDLs can be separated by ion-exchange chromatography into a-HDL and b-HDL. Although they have similar chemical compositions (Bernardi and Cook, 1960), notable differences were observed in sialic acid and in their protein-bound phosphorus content (0.39 and 0.19% P, respectively) (Kurisaki et al., 1981). Each monomer of HDL consists of five major polypeptides with molecular weight ranging from 32 to 105 kDa, this latter being the main one. Plasma proteins in egg yolk Lipovitellenin (LDL) LDL is the main constituent of yolk (2/3 of the total yolk dry matter) and is mainly found in plasma but a residual portion is included in granules. The LDL structure was described as spherical particles with a lipid core surrounded by a layer of phospholipids and proteins (Evans et al., 1973). LDL is composed of 11–17% proteins and 83–89% lipids. There is little agreement among the various molecular weight values of the apoproteins of LDL published to date. Polypeptides reported range between 15 and 180 kDa (Anton et al., 2003; Yamauchi et al., 1976) and higher molecular weights up to 225–240 kDa have also been found (Itoh et al., 1986; Mine, 1998). Owing to its low density (0.98), LDL is soluble in aqueous solution whatever the pH and ionic conditions. Livetins Livetins comprise the water-soluble globular protein fraction which is composed of a-, b-, and g-livetins in the ratio 2:5:3, respectively, in the yolk (Bernardi and Cook, 1960). The molecular weights of a-, b-, g-livetins are 80, 45 and 150 kDa, respectively (Martin et al., 1957). The a- and b-livetins are thermolabile, whereas the g-livetins are more thermostable.
19.3 Biological activities of egg proteins and peptides and applications in the nutraceutical industry Biologically-active peptides can be produced in vitro through enzymatic hydrolysis of egg proteins. Most bioactive peptides are produced using gastrointestinal enzymes, such as pepsin, chymotrypsin and trypsin. For example, angiotensin-converting enzyme (ACE) inhibitory peptides from ovalbumin are usually produced by peptic hydrolysis. In general, they contain 2–20 amino acid residues per molecule but some may consist of more than 20 amino acids. These peptides can exhibit various activities including ACE-inhibitory, antihypertensive, antimicrobial, vasorelaxant and
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602 Separation, extraction and concentration processes immunomodulant actions. Some bioactive peptides can exert more than one activity (Table 19.4). 19.3.1 ACE inhibitor, antihypertensive and vasorelaxing activities Various peptides derived from proteolytic digestion of egg white proteins contain a number of peptide sequences having in vitro ACE-inhibitory activity. In vivo studies showed some antihypertensive and/or vasorelaxing activities in spontaneously hypertensive (SHR) rats models. Fujita et al. (1995) reported the first bioactive peptide, named ovokinin (OVA 358–365), derived from peptic digestion of ovalbumin. Ovokinin showed relaxing activity in canine mesenteric artery and possessed a high ACE-inhibitory activity, IC50 = 3.2 mM (Miguel et al., 2004). A hexapeptide corresponding to the 2–7 fragment of ovokinin (RADHPF) was isolated from a chymotryptic digest of ovalbumin. Ovokinin (2–7) showed lower ACEinhibitory activity (IC50 > 400 mM) but exerted a potent vasorelaxant and antihypertensive effect upon oral administration to SHR rats at a dosage of 10 mg kg–1 (Matoba et al., 1999). Non-hydrolyzed whole egg white did not have ACE-inhibitory property whereas its digestion with pepsin resulted in high ACE-inhibitory activity with IC50 = 55.3 mg mL–1 after hydrolysis for 3 h. Treatment with trypsin and chymotrypsin induced very poor ACE-inhibitory activity values, indicating the importance of the enzyme specificity in producing peptides with ACEinhibitory activities (Miguel et al., 2004). Furthermore, these workers showed that the active peptides could be enriched from the hydrolysate using ultrafiltration with a 3 kDa molecularweight-cut-off (MWCO) membrane. The permeate generated contained peptides with molecular mass <3 kDa that presented 10 times more ACE-inhibitory activity than the retentate (IC50 = 34.5 mg mL–1 and IC50 = 298.4 mg mL–1, respectively) and were identified as RADHPFL and YAEERYPIL. These two peptides exhibited a significant antihypertensive effect in SHR rats at a dose of around 2 mg kg–1 (Miguel et al., 2005). Proteolytic degradation with pepsin and pancreatic extract of RADHPFL and YAEERYPIL released two main fragments RADHP and YPI which decreased blood pressure at doses of 2 mg kg–1 after 2 h of administration but did not exhibit ACE-inhibitory activity (Miguel et al., 2006). Vasorelaxing activity has also been found in the sequences RADHP and YPI and show common features with other vascular relaxing peptides. Peptides with antihypertensive activity have also been generated by enzymatic hydrolysis of egg yolk. Oligopeptides of 1 kDa and less showed ACE-inhibitory and suppressed the development of hypertension in SHR rats in a dose-dependent manner after oral administration for 12 weeks (Yoshii et al., 2001).
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Table 19.4 Bioactive peptides derived from egg proteins Bioactive peptide
Fragment
Sequence
Enzyme
Bioactivity
Ovalbumin Ovalbumin Ovalbumin Ovalbumin Ovalbumin Lysozyme Lysozyme Ovalbumin Ovalbumin Ovalbumin Ovalbumin Ovalbumin Ovalbumin Ovalbumin Ovalbumin Lysozyme Phosvitin
Ovokinin Ovokinin (2–7) Unnamed Unnamed Unnamed Unnamed Unnamed Unnamed Unnamed Unnamed Unnamed Unnamed Unnamed Unnamed Unnamed Unnamed PPP
358–365 359–364 359–365 106–114 36–40 98–108 15–21 36–40 36–42 111–119 143–148 159–165 127–138 155–159 268–276 98–112 Unknown
FRADHPFL RADHPF RADHPFL YAEERYPIL SALAM IVSDGDGMNAW HGLDNYR SALAM SALAMVY YPILPEYLQ ELINSW NVLQPSS AEERYPILPEYL GIIRN TSSNVMEER IVSDGNGMNAWVAWR Unknown
Pepsin Chymotrypsin Pepsin Pepsin Pepsin Pepsin and trypsin Pepsin and trypsin Trypsin Trypsin Trypsin Trypsin Trypsin Chymotrypsin Chymotrypsin Chymotrypsin Clostripain Trypsin
Vasorelaxing/antihypertensive Vasorelaxing/antihypertensive ACE/antihypertensive ACE/antihypertensive ACE/antimicrobial Antimicrobial Antimicrobial Antimicrobial Antimicrobial Antimicrobial Antimicrobial Antimicrobial Antimicrobial Antimicrobial Antimicrobial Antimicrobial Antioxidant/anti-inflammatory
PPP, phosphopeptides.
Fractionation of egg proteins and peptides 603
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Protein precursor
604 Separation, extraction and concentration processes 19.3.2 Antimicrobial activity Lysozyme is an enzyme found in many cells, tissues and secretion of organisms (Jollès and Jollès, 1984). Lysozyme extracted from hen’s egg white exerts bacteriolytic activity on bacterial cells walls. It is most effective against some specific Gram-positive bacteria and, to a lesser extent, against Gramnegative bacteria owing to their composition of peptidoglycans (Pellegrini et al., 1992). Various strategies have been used to increase antimicrobial activity of lysozyme against Gram-negative bacteria. A number of chemical modifications aimed at inserting a hydrophobic moiety (Ibrahim et al., 1991; 1993), adding hydrophobic peptides to the C-terminus of lysozyme (Ibrahim et al., 1992; 1994) or lypophilizing of lysozyme by a fatty acid having a different chain length (Ibrahim et al., 1993; Liu et al., 2000). Nakamura et al. (1991, 1996) and Nakamura and Kato (2000) enhanced antimicrobial activity against Gram-negative bacteria by conjugating lysozyme with polysaccharides. Proteolytic digestion of lysozyme with clostripain has been reported to yield an antimicrobial peptide fragment (f98–112) (Ibrahim et al., 2001; Pellegrini et al., 1997). This peptide fragment has 15 amino acids and is located close to the C-terminal end of lysozyme; it possesses antibacterial activity against both Gram-positive and Gram-negative bacteria (Pellegrini et al., 1997). From a structural point of view, this sequence is a part of a helix-loop-helix domain located at the upper lip of the active site cleft of the lysozyme, which occurs in several bactericidal peptides (Ibrahim et al., 2001; Pellegrini, 2003). Chemical substitution of some amino acids to confer a net positive charge of the peptide 98–112 allowed further enhancement of these bactericidal properties. Hydrophobicity was also identified as an important factor in promoting interactions with bacterial membranes and had a strong impact on the bactericidal activity of this pentadecapeptide (Ibrahim et al., 2001; Pellegrini, 2003). Mine et al. (2004) characterized a novel antimicrobial peptide from chicken egg white lysozyme obtained by peptic and tryptic digestion. The peptide (98–108), located in the helixloop-helix domain and the peptide (15–21) possesses antimicrobial activity against Escherichia coli and Staphylococcus aureus. Commercial peptide mixtures from hen egg lysozyme, produced by partial enzymatic hydrolysis of lysozyme with pepsin, also exhibit a powerful antimicrobial activity to control Bacillus species in food products (Abdou et al., 2007). Ovalbumin contains bactericidal peptides that can be released by enzymatic digestion. Five antimicrobial peptides (SALAM, SALAMVY, YPILPEYLQ, ELINSW and NVLQPSS) from tryptic digests and peptides AEERYPILPEYL, GIIRN and TSSNVMEER from chymotryptic digests were found to be strongly active against Bacillus subtilus and to a lesser extent against the Gram-positive and Gram-negative bacteria studied (Pellegrini et al., 2004). Ovotransferrin displays in vivo and in vitro antibacterial action owing to both its iron-binding activity and its 92 residues N-terminal domain (Ibrahim et al., 1998; 2000). The cationic peptide of hen ovotransferrin, called OTAP
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Fractionation of egg proteins and peptides 605 92, corresponds to the sequence of ovotransferrin located at the lip of the iron-binding cleft of the N lobe of ovotransferrin. This peptide exhibits bactericidal activity against Gram-positive and Gram-negative bacteria by penetrating the bacterial outer membrane by self-promoted uptake and causing damage to the cytoplasmic membranes (Ibrahim et al., 2000). 19.3.3 Antioxidant activity Numerous phosphoserine residues grouped in the amino acids sequence of phosvitin, are responsible for the iron-binding capacity and they do confer to phosvitin a potential antioxidant activity. The metal-binding capacity of phosvitin can control iron-catalyzed or copper-catalyzed lipid oxidation to a Fe2+: phosvitin ratio of 30:1 and up to a Cu2+: phosvitin ratio of 1:1. Moreover, pasteurization does not affect the iron-binding capacity and antioxidant potential of phosvitin (Lu and Baker, 1986). Nakamura et al. (1998) produced a novel macromolecular antioxidant by conjugating phosvitin with galactomannan which could withstand a sterilization treatment at 121 °C for 15 min. Peptides generated from the digestion of various proteins are reported to have antioxidant activities. Jiang and Mine (2000) have developed phosvitin phosphopeptides (PPP) with molecular weight values of 1–3 kDa from tryptic hydrolysis following partial alkaline dephosphorylation. These peptides showed antioxidant activity against oxidative stress in human intestinal epithelial cells in an in vitro assay using Caco-2 cells (Katayama et al., 2006). In another study, PPP can up-regulate cellular glutathione biosynthesis associated enzymes activity and antioxidant activities in oxidative stress induced intestinal epithelial cells (Katayama et al., 2007). Although PPP have not yet been characterized, the antioxidant activity of PPP may be associated with molecular weight, phosphorus content, hydrophobicity and/ or amino acid composition. Ovalbumin was found to possess a strong antioxidant activity against linolenic acid and docosahexaenoic acid (Nara et al., 1995). Three short peptides from ovalbumin protein hydrolysates have been identified and it has been suggested that metal chelation plays an important role in their antioxidant activity (Hamachio and Hasegawa, 1989).
19.4 Available technologies for the fractionation of egg proteins and peptides, and applications in the nutraceutical industry Egg proteins are becoming important in the poultry product industry because of their technological and functional properties. Moreover, hen egg possesses many biologically active proteins that may have numerous applications in the food and pharmaceutical industries (Mine, 2007). Accordingly, there
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606 Separation, extraction and concentration processes is a need to develop efficient, simple and cost-effective methodologies for isolation and purification of egg proteins and their peptides. Figure 19.1 summarizes the technological options available for fractionating egg proteins and their peptides. The contemporary industrial processes comprise three main approaches alone or in combination, namely precipitation, chromatography and membrane processes. The appropriate selection of a given technological approach has to take into account some key physicochemical characteristics of the proteins/peptides to be separated from complex mixtures. Table 19.5 summarizes some of the structural properties of bioactive peptides derived from egg protein. These are often related to the amino acid sequence of the peptides. For example, ACE-inhibitory peptides commonly contain a positive side-chain charge on the C-terminal residue. This feature can be exploited by using ion-exchange chromatography and/ or membrane separation with charged polymeric material. Isoelectric precipitation Precipitation
Salting-out Organic solvents Gel permeation
Chromatography
Ion-exchange Reverse-phase/hydrophobic Microfiltration
Membrane separations
Ultrafiltration Nanofiltration
Fig. 19.1 Technological approaches and processes to fractionate egg proteins/ peptides. Table 19.5 Common structural properties of bioactive peptides Activity
Structural element
Antimicrobial
Short sequences (<40 residues) Helix-loop-helix motif Positive charge and hydrophobic properties Rich in SerP High amounts of His and hydrophobic amino acids Molecular size Positive charge in C-terminal residue Presence of aromatic and/or hydrophobic amino acids in C-terminal residue Arg and Tyr residues in N-terminal position
Antioxidant ACE inhibitory Vasorelaxing
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Fractionation of egg proteins and peptides 607 19.4.1 Precipitation Precipitation of egg proteins can be achieved by changes in the solvent properties including pH and/or ionic strength modifications and by addition of organic solvents. The most commonly used precipitating agents are ammonium sulfate, acetone or ethanol. Precipitation by ionic strength and/or pH modifications Purification of ovalbumin, the main egg-white protein, involves its precipitation at specific salt concentration, pH and temperature values. Ammonium sulfate or sodium sulfate are used for ovalbumin precipitation. Lysozyme is the main egg-white protein extracted on an industrial scale for commercial applications. Lysozyme can be precipitated from egg-white proteins by increasing the pH to 9.5 and by addition of sodium chloride to a final concentration of 5% (w/v) (Alderton and Fevold, 1946). Ovotransferrin has been purified with ammonium sulfate in crystalline form, both as the iron complex and as the iron free protein (Warner and Weber, 1951). Ovomucin can be precipitated by simple dilution with two or three volumes of water and by lowering pH to its pI of 4.5–5.0. A washing step with salt solution is required to dissociate ovomucin from lysozyme and finally salts are washed out with water (Cotterill and Winter, 1955). Purification of the immunoglobulin Y (IgY) involves removal of lipids and lipoproteins from egg yolk. Various strategies involving detergents such as SDS (Sriram and Ygeeswaran, 1999) and polysaccharides (Hatta et al., 1988; 1990), solvents (Sriram and Ygeeswaran, 1999) and polyethyleneglycol (Akita and Nakai, 1993) were used to remove lipids from egg yolk. Salt precipitation can thereafter be achieved for IgY purification. Ammonium sulfate and sodium sulfate were used and concentration levels were dependent on the IgY yield and purity (Akita and Nakai, 1992). Although purification by precipitation of proteins for preparative purposes can be achieved, these methods suffer from major deficiencies. Precipitation by salts leads to a protein extract and to by-products having a high salt concentration. Several solubilizations and crystallization cycles must be performed to obtain highly purified proteins. Precipitation by organic solvents Addition of an organic solvent to protein solution results in important modifications of the dielectric constant of the medium and to the weakening of protein interactions with water. Organic solvents have an affinity for the hydrophobic surfaces of the proteins and this result in denaturation of the proteins along with precipitation. However, low temperatures (–5 to 0 °C) are necessary to minimize protein denaturation. Ovomucoid is the main egg white protein isolated by using organic solvent. First, precipitation of other egg white proteins is eliminated with 2.7% trichloroacetic acid at pH 3.5. Then, ovomucoid is precipitated by addition of acetone (Lineweaver and Murray, 1947).
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608 Separation, extraction and concentration processes After salt precipitation, using solvent such as cryoethanol gave an IgY isolate of 93% purity (Akita and Nakai, 1992); ethanol is used in large-scale IgG production (Horikoshi et al., 1993). Precipitation by organic solvents is widely used in the chemical industry. However, the necessity of several precipitation steps to obtain a purified protein make this process extremely difficult to scale up and the low temperature restricts the scale of production. Moreover, the use of organic solvent is considered undesirable in the food and nutraceutical industrial environment. 19.4.2 Chromatographic methods Egg proteins are characterized by specific physicochemical characteristics such as their size, pI, amino acid composition, ability to bind metal ions (ovotransferrin and lysozyme), protein–protein interaction potential (lysozyme and ovomucin) or ligand-binding capacity (avidin and flavoprotein). Based on these characteristics, several methods, including gel-permeation, ionexchange and reversed-phase chromatography, were used for the separation or purification of egg white and yolk proteins. Gel-permeation chromatography Gel-permeation chromatography (GPC) separates molecules according to the difference in their size. This technique is based on the penetration of molecules into the cavities of a macroporous support, mostly made from hydrophilic gels of dextran, agarose or polyacrylamide. In general, molecules with a hydrodynamic diameter smaller than the diameter of the pores in the support diffuse into the matrix, whereas molecules with larger diameters are excluded, thus passing through more quickly. Several proteins of egg-white and yolk proteins have been isolated by GPC. Purification of ovomucin from egg-white proteins was carried out using a Sepharose 4B and Superose 6 preparative-grade columns. Gel permeation on Superose 6 permitted simultaneous purification of ovomucin and lysozyme, but with Sepharose 4B another compound eluted with the ovomucin and this may be ovostatin (Awade et al., 1994; Young and Gardner, 1972). GPC was used in combination with other chromatographic procedures. After gel permeation on Superose 6, ovotransferrin and ovalbumin were isolated by anion-exchange chromatography on Q Sepharose Fast Flow. The two-step purification procedure gave 80, 100, 80 and 100% purity for ovomucin, lysozyme, ovotransferrin and ovalbumin, respectively. In another study, a better separation of ovomucin and lysozyme was obtained with a Superose 12 HR 10/30 column and anion-exchange or reversed-phase high-performance liquid chromatography (RP-HPLC) were proposed to purify other egg white proteins (Awade and Efstathiou, 1999). GPC was also used to isolate phosvitin, the very low-density lipoproteins from delipidated egg yolk proteins (Tsutsui and Obara, 1984). Abe et al.
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Fractionation of egg proteins and peptides 609 (1982) separated phosvitin into a- and b-phosvitin by gel permeation on a Sephadex G-200 column. These two components differ in their amino acid composition, their carbohydrate content and their precipitation in the presence of calcium (Itoh et al., 1983). GPC is most useful in the separation of egg proteins owing to their wide range of molecular mass distribution. However, GPC allows compounds having a molecular size between two relatively close thresholds to be separated. Because of the fragile nature of many soft gels used for GPC and because interaction with proteins reduces the potential utilization of GPC for egg protein fractionation, large-scale processes may not be easily applicable. Ion exchange chromatography Ion-exchange chromatography (IEC) is the most widely used large-scale method for the purification of proteins and other charged molecules. In cation-exchange chromatography, positively charged molecules are attracted to a negatively charged solid support. Conversely, in anion-exchange chromatography, negatively charged molecules are attracted to a positively charged solid support. A variety of resins may be used to fractionate lysozyme from egg white proteins. Among weakly acidic cation-exchange resins, Duolite C-464 exhibited an efficient method for separation and recovery of active lysozyme (86%). However, a significant amount of avidin co-eluted with lysozyme (Durance and Nakai, 1988; Li-Chan et al., 1986). Roy et al. (2003) described an integrated easily scalable process that is used to simultaneously purify the major egg white proteins. This technique involves a cation exchanger, Streamline™ SP and differential precipitation, followed by dye–ligand chromatography. Purified proteins were obtained and the yields of lysozyme, ovomucoid and ovalbumin were 77, 94 and 98%, respectively. Guérin-Dubiard et al. (2005) developed a procedure for fractionating the whole egg white using three successive chromatography steps. High-purity levels of lysozyme (95%), ovotransferrin (89%), ovalbumin (91%) and flavoprotein (100%) were obtained. Connelly and Taborsky (1961) separated phosvitin by use of stepwise salt elution on a DEAE cellulose column into two sub-fractions: a major fraction called 0.30, and a minor one called 0.35 phosvitin. The fractions had approximately the same amino acid composition but differed in their metal content and chemical stability at alkaline pH. Recently, Castellani et al. (2003) described a new purification method that includes a first extraction step based on insolubility of Mg2+/phosvitin salts and a second step by ion-exchange chromatographic fractionation avoiding organic solvents. Purification gives a- and b-phosvitin that are free from contaminants, highly purified (>98%) and metal-free phosvitin. IEC is suitable for fractionating egg proteins with high purity and, until now, the application of IEC on an industrial scale concerned the major egg white proteins.
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610 Separation, extraction and concentration processes Reversed-phase high-performance liquid chromatography The separation mechanism in RP-HPLC is based on the hydrophobic interaction of proteins/peptides with a non-polar stationary phase. Itoh et al. (1991) investigated the recovery of hydrophobic egg white proteins using RP-HPLC at room temperature. Under conventional gradient elution conditions, the recovery of ovalbumin was poor. However, under fast separation conditions, the recovery of the proteins was dramatically improved. In contrast, under both chromatographic conditions approximately 100% of lysozyme was recovered. For RP-HPLC of lipid-free egg yolk proteins solvents containing formic acid were used. Although several peaks were unidentified, three groups (granules proteins, low-density lipoprotein apoproteins, and livetins) were separated (Sheumack and Burley, 1988). RP-HPLC is not used to fractionate egg white and yolk proteins as much as other chromatographic methods. A gradient of a mixture of a non-polar solvent is required for elution and protein denaturation may occur during RP-HPLC process. 19.4.3 Membrane processes Pressure-driven membrane-based separation processes comprise microfiltration (MF), ultrafiltration (UF), nanofiltration (NF) and reverse osmosis (RO) and are used for protein separation/purification (Cheryan, 1998). Depending on MWCO, various membranes can be used for the separation of different molecular weight proteins. MF membranes are suitable for the separation of particles in the size range of 0.1–10 mm whereas UF membranes are used for macromolecules with a molecule weight of 1–300 kDa. Compared with UF, NF membranes have a smaller pore size and retain smaller organic molecules, 200–2000 Da. RO drives water molecules from a low to a high concentration region and results in the concentration of salt molecules on the retentate side of the membrane. Size is the main sieving mechanism in MF and RO whereas the separation mechanism is normally explained in terms of charge and/or size effects in UF and NF. UF is mainly used for protein concentration, desalting, clarification and protein fractionation. Protein fractionation is strongly influenced by operating and physicochemical parameters such as pH and ionic strength that affect the protein–protein and protein–membrane interactions and thus the selectivity. Model egg white protein solutions were fractionated with modified and unmodified 50 kDa MWCO polysulfone UF membranes (Eshani et al., 1997). Electrostatic exclusion prevented ovotransferrin and lysozyme from permeating the membrane and this led to a permeate containing almost pure ovalbumin. Fractionation was obtained at pH 4.8 without salts on both UF membranes. In another study, UF polysulfone membrane pre-treated with myoglobin enhanced lysozyme purification from chicken egg white compared with that obtained with the native membrane (Ghosh and Cui, 2000a). Ghosh and Cui (2000b) reported the effect of pH on fractionation of
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Fractionation of egg proteins and peptides 611 lysozyme from chicken egg white. A high transmission of lysozyme was observed with increasing pH showing the influence of pH on the purification of lysozyme. In some instances, a combination of UF membranes is useful for protein fractionation. A two-step ultrafiltration process using 50 and 25 kDa MWCO membranes was achieved to purify lysozyme. In a first step, lysozyme was enriched in the UF permeate of 50 kDa MWCO. Then, the 25 kDa MWCO UF membrane was used to purify the lysozyme. This combination gave a high productivity and purity for the separation process. Similarly, a two-stage UF technique was used to separate ovalbumin from chicken egg white. In the first stage, ovalbumin was retained by a 30 kDa flat disk polyethersulfone membrane; this retentate was fractionated using 50 kDa flat disk polyethersulfone membrane in the second stage. A high-purity ovalbumin (98.7%) could be produced using a two-step UF process under the experimental conditions studied (Datta et al., 2009). Separation of lysozyme from chicken egg white has been investigated by using 30 kDa hollow fiber polysulfone membranes. Higher lysozyme transmission (>90%) was observed under optimized conditions and moderately pure lysozyme (80–90%) could be obtained by carrying out a diafiltration step (Ghosh et al., 2000). Production of ACE-inhibitory peptides from egg white proteins with various proteolytic enzymes was carried out using a membrane reactor. Egg white proteins hydrolyzed with thermolysin produced the highest ACE-inhibitory potential (IC50 = 54.1 mg mL–1). Thermolysin hydrolysate was fractionated using UF membrane of MWCO of 10, 3 and 1 kDa. A lower IC50 (17.2 mg mL–1) was obtained in the UF permeate with 1 kDa MWCO (Chiang et al., 2008). By a modification of polymeric forms of lysozyme under optimal conditions of UF, a preparation comprising 53.3% of lysozyme polymeric forms, i.e. 33.2% of dimer and 20.1% of trimer, was obtained. Modified lysozyme preparation by UF showed the highest bacteriostatic activity against Gramnegative bacteria selected (Lesnierowski et al., 2009). A water-soluble plasma protein from egg yolk granules was obtained with various simple water dilutions, followed by centrifugation or filtration. Two factors were critical, pH and egg yolk dilution, for IgY recovery. Optimum recovery of IgY (93–96%) was obtained by a six-fold water dilution at pH 5.0–5.2 with incubation for 6 h at 4 °C. A 100 kDa MWCO extracted IgY from water-soluble plasma upon the separation of crude IgY by ammonium sulfate precipitation. A maximal recovery of IgY (> 98%) was obtained and this value was similar to the degree of purity obtained by gel filtration (Akita and Nakai, 1992). A serial filtration approach, involving dilutions, paper filtration and delipidation using hydrophobic filters or using different UF membranes, was developed by Kim and Nakai (1996; 1998). The delipidated water-soluble fraction was thereafter purified using 100 kDa UF membranes and both high recoveries (72–89%) and purity (74–99%) were obtained. Although the fractionation of proteins through a membrane depends
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612 Separation, extraction and concentration processes on physicochemical parameters such as pH and salt concentration (Kim and Nakai, 1996; 1998), the surface properties of the membrane are also important for the successful protein fractionation. Hernandez-Campos et al. (2010) purified IgY from egg yolk water-soluble protein by UF–diafiltration with different membranes. The best selectivity and purification factors were obtained without salt at pH 5.7 and 6.7 using polyethersulfone and modified polyethersulfone of 100 MWCO membrane, respectively. Compared with chromatographic methods, membrane separation techniques offer the advantage of lower cost and ease of scale-up for commercial production. Membrane-based separation techniques are powerful tools for fractionation/purification of proteins and bioactive peptides. UF membranes can be used as preliminary step for the removal of enzymes and non-hydrolyzed proteins, and to further fractionate the peptide mixture. Various combinations of precipitation and UF using selective membranes can be applied for the fractionation of food proteins. In addition to selective membranes, chromatographic methods can be employed successfully to isolate specific bioactive compounds from a complex mixture. UF membranes were successfully used to enrich peptide fractions. Membranes consisting of negatively charged materials such as NF were used to desalt or to fractionate acidic peptides from hydrolysate mixture (Chay Pak Ting et al., 2007; Wijers et al., 1998). 19.4.4 Other separation methods A new isolation method was described for IgY. Use of food-grade products such as sodium alginate or l-carrageenan was investigated for precipitation of yolk lipoproteins from the water-soluble fraction to purify IgY. Polysaccharides exhibit satisfactory yields of IgY (Hatta et al., 1988; 1990). In another study, the optimal separation of IgY was achieved with 0.15% pectin and at pH 5.0, implying that the interactions between polysaccharides and lipoproteins were mainly ionic bonds, hydrophobic interactions and hydrogen bonds (Chang et al., 2000).
19.5 Conclusion and perspectives Considering that the discovery of novel bioactive peptides and their possible functions and health benefits are constantly increasing, there is a growing need to develop new technologies for the production of specific bioactive peptides from hydrolysate mixtures. Membrane-based separation techniques have traditionally been used to separate molecules of different sizes and lowmolecular-weight components from proteins. UF, alone and in combination with other techniques such as NF and chromatography has shown its potential to separate peptide mixtures on a large scale. However, the further research into membrane areas (such as membrane materials and chemistry, © Woodhead Publishing Limited, 2010
Fractionation of egg proteins and peptides 613 module configurations) for the complete separation/purification of proteins/ peptides is still needed. The addition of affinity ligands to chelate proteins/ peptides could be useful in obtaining maximal selectivity by combining with membrane processes.
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614 Separation, extraction and concentration processes Connelly C and Taborsky G (1961), Chromatographic fractionation of phosvitin, J Biol Chem, 236, 1364–1368. Cotterill O J and Winter A R (1955), Egg white lysozyme. 3. The effect of pH on the lysozyme-ovomucin interaction, Poult Sci, 34, 679–686. Cunningham F E and Lineweaver H (1965), Stabilization of egg white proteins to pasteurization temperature above 60 °C, Food Technol, 19, 136–141. Datta D, Bhattacharjee S, Nath A, Das R, Bhattacharjee C and Datta S (2009), Separation of ovalbumin from chicken egg white using two-stage ultrafiltration technique, 66, 353–361. Donovan J W and Mapes C J (1976), A differential scanning calorimetric study of conversion of ovalbumin to S-ovalbumin in eggs, J Sci Food Agric, 27, 197–204. Durance T D and Nakai S (1988), Simultaneous isolation of avidin and lysozyme from egg albumen, J Food Sci, 53, 1096–1102. Eshani N, Parkkinen S and Nyström M (1997), Fractionation of natural and model eggwhite protein solutions with modified and unmodified polysulfone UF membranes, J Membr Sci, 123, 105–119. Evans R J, Bauer D H, Bandemer S L, Vaghefi S B and Flegal C J (1973), Structure of egg yolk very low density lipoprotein. Polydispersity of the very low density lipoprotein and the role of lipovitellenin in the structure. Arch Biochem Biophys, 154, 493–500. Fujita H, Usui H, Kurahashi K and Yoshikawa M (1995), Isolation and characterization of ovokinin a bradykinin B1 agonist peptide derived from ovalbumin, Peptides, 16, 785–790. Ghosh R and Cui Z F (2000a), Protein purification by ultrafiltration with pre-treated membrane, J Membr Sci, 167, 47–53. Ghosh R and Cui Z F (2000b), Purification of lysozyme using ultrafiltration, Biotechnol Bioeng, 68, 191–203. Ghosh R, Silva S S and Cui Z F (2000), Lysozyme separation by hollow fibre ultrafiltration, Biochem Eng J, 6, 19–24. Greengard O, Sentenac A and Mendelsohn N (1964), Phosvitin, the iron carrier of egg yolk, Biochim Biophys Acta, 90, 406–407. Guérin-Dubiard C, Pasco M, Hietanen A, Quiros del Bosque A, Nau F and Croguennec T (2005), Hen egg white fractionation by ion-exchange chromatography, J Chromatogr A, 1090, 58–67. Hamachio Y and Hasegawa M (1989), EPA powder as a functional ingredient, New Food Ind, 31, 12–16. Hatta H, Kim M and Yamamoto T (1990), A novel isolation method for hen egg yolk antibody, ‘IgY’, Agric Biol Chem, 54, 2531–2535. Hatta H, Sim J S and Nakai S (1988), Separation of phospholipids from egg yolk and recovery of water-soluble proteins, J Food Sci, 53, 425–427, 431. Hernandez-Campos F J, Brito-De La Fuente E and Torrestiana-Sanchez B (2010), Purification of egg yolk immunoglobulin (IgY) by ultrafiltration: effect of pH, ionic strength and membrane properties, J Agric Food Chem, 58(1), 187–193. Horikoshi T, Hiraoka J, Saito M and Hamada S (1993), IgG antibody from hen egg yolks: purification by ethanol fractionation, J Food Sci, 58, 739–742. Ibrahim H R, Iwamori E, Sugimoto Y and Aoki T (1998), Identification of a distinct antibacterial domain within the N-lobe of ovotransferrin, Biochim Biophys Acta, 1401, 289–303. Ibrahim H R, Kato A and Kobayashi K (1991), Antimicrobial effects of lysozyme against Gram-negative bacteria due to covalent binding of palmitic acid, J Agric Food Chem, 39, 2077–2082. Ibrahim H R, Kobayashi K and Kato A (1993), Length of hydrocarbon chain and antimicrobial action to Gram-negative bacteria of fatty acylated lysozyme, J Agric Food Chem, 41, 1164–1168.
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Fractionation of egg proteins and peptides 615 Ibrahim H R, Sugimoto Y and Aoki T (2000), Ovotransferrin antimicrobial peptide (OTAP 92) kills bacteria through a membrane damage mechanism, Biochim Biophys Acta, 1523, 196–205. Ibrahim H R, Thomas U and Pellegrini A (2001), A helix-loop-helix peptide at the upper lip of the active site cleft of lysozyme confers potent antimicrobial activity with membrane permeabilization action, J Biol Chem, 276, 43767–43774. Ibrahim H R, Yamada M, Kobayashi K and Kato A (1992), Bactericidal action of lysozyme against Gram-negative bacteria due to insertion of a hydrophobic pentapeptide into its C-terminus, Biosci Biotechnol Biochem, 56, 1361–1363. Ibrahim H R, Yamada M, Matsushita M, Kobayashi K and Kato A (1994), Enhanced bactericidal action of lysozyme to Escherichia coli by inserting a hydrophobic pentapeptide into its C-terminus, J Biol Chem, 18, 5059–5063. Ishihara H, Takahasi N, Ito J, Takeuchi E and Tejima S (1981), Either high-manonose-type or hybrid-type oligosaccharide is linked to the same asparagine residue in ovalbumin, Biochim Biophys Acta, 669, 216–221. Itoh T, Abe Y and Adachi S (1983), Comparative studies on the a- and b-phosvitin from hen’s egg yolk, J Food Sci, 48, 1755–1757. Itoh T, Kubo M and Adashi S (1986), Isolation and characterization of major apoproteins from hen’s egg yolk granule, J Food Sci, 51, 1115–1117. Itoh T, Miyakazi J, Sugawara H and Adachi S (1987), Studies on the characterization of ovomucin and chalaza of the hen’s egg, J Food Sci, 52, 1518–1521. Itoh H, Nimura N, Kinoshita T, Nagae N and Nomura M (1991), Fast protein separation by reversed-phase high-performance liquid chromatography on octadecylsilyl-bonded nonporous silica gel. II. Improvement in recovery of hydrophobic proteins, Anal Biochem, 199, 7–10. Jiang B and Mine Y (2000), Preparation of novel functional oligophosphopeptides from hen egg yolk phosvitin, J Agric Food Chem, 48, 990–994. Jollès P and Jollès J (1984), What’s new in lysozyme research?, Mol Cell Biochem, 63, 165–189. Katayama S, Ishikawa S I, Fan M Z and Mine Y (2007), Oligophosphopeptides derived from hen egg yolk phosvitin up regulate g-glutamylcysteine synthetase and antioxidant enzymes against oxidative stress in Caco-2 cells, J Agric Food Chem, 55, 2829–2835. Katayama S, Xu X, Fan M Z and Mine Y (2006), Antioxidant stress activity of oligophosphopeptides derived from hen egg yolk phosvitin in Caco-2 cells, J Agric Food Chem, 54, 773–778. Kato I, Schrode J, Kohr W J and Laskowski Jr M (1987), Chicken ovomucoid: determination of its amino acid sequence, determination of trypsin reactive site, and preparation of all three of its domains, Biochemistry, 26, 193–201. Kim H and Nakai S (1996), Immunoglobulins separation from egg yolk: a serial filtration system, J Food Sci, 61, 510–513. Kim H and Nakai S (1998), Simple separation of immunoglobulin from egg yolk by ultrafiltration, J Food Sci, 63, 485–490. Kurisaki J K, Yamauchi H, Ishiki H and Ogiwara S (1981), Differences between a- and b-lipovitellin from hen egg yolk, Agric Biol Chem, 45, 699–704. Lesnierowski G, Kijowski J and Cegielska-Radziejewska R (2009), Ultrafiltrationmodified chicken egg white lysozyme and its antibacterial action, Int J Food Sci Technol, 44, 305–311. Li-Chan E, Nakai S, Sim J, Bragg D B and Lo K V (1986), Lysozyme separation from egg white by cation exchange column chromatography, J Food Sci, 51, 1032–1036. Li-Chan E and Nakai S (1989), Biochemical basis for the properties of egg white, Crit Rev Poult Biol, 2, 21–57. Lineweaver H and Murray C W (1947), Identification of the trypsin inhibitor of egg white with ovomucoid, J Biol Chem, 171, 565–581.
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616 Separation, extraction and concentration processes Liu S T, Sugimoto T, Azakami H and Kato A (2000), Lipophilization of lysozyme by short and middle chain fatty acids, J Agric Food Chem, 48, 265–269. Losso J N, Bogumil R and Nakai S (1993), Comparative studies of phosvitin from chicken and salmon egg yolk, Comp Biochem Physiol, 106, 919–923. Lu C L and Baker R (1986), Characteristics of egg yolk phosvitin as an antioxidant for inhibiting metal-catalyzed phospholipid oxidations, Poult Sci, 65, 2065– 2070. Martin W G, Vandegaer J E and Cook W H (1957), Fractionation of livetin and the molecular weights of the a- and b-components, Can J Biochem Physiol, 35, 241–250. Matoba N, Usui H, Fujita H and Yoshikawa M (1999), A novel anti-hypertensive peptide derived from ovalbumin induces nitric oxide-mediated vasorelaxation in an isolated SHR mesenteric artery, FEBS Lett, 452, 181–184. Matsuda T, Watanabe K and Nakamura R (1982), Immunochemical studies on thermal denaturation of ovomucoid, Biochim Biophys Acta, 707, 121–128. Mecham D K and Olcott H S (1949), Phosvitin, the principal phosphoprotein of egg yolk, J Am Chem Soc, 71, 3670–3679. Miguel M, Aleixandre A, Ramos M and Lopez-Fandino R (2006), Effect of simulated gastrointestinal digestion on the antihypertensive properties of ACE-inhibitory peptides derived from ovalbumin, J Agric Food Chem, 54, 726–731. Miguel M, Lopez-Fandino R, Ramos M and Aleixandre A (2005), Short-term effect of egg-white hydrolysate products on the arterial blood pressure of hypertensive rats, Br J Nutr, 94, 731–737. Miguel M, Recio I, Gomez-Ruiz J A, Ramos M and Lopez-Fandino R (2004), Angiotensin I-converting enzyme inhibitory activity of peptides derived from egg white proteins by enzymatic hydrolysis, J Food Prot, 67, 1914–1920. Mine Y (1998), Adsorption behaviour of egg yolk low-density lipoproteins in oil-in-water emulsions, J Agric Food Chem, 46, 36–41. Mine Y (2007), Egg proteins and peptides in human health-chemistry, bioactivity and production, Curr Pharm Des, 13, 875–884. Mine Y, Ma F and Lauriau S (2004), Antimicrobial peptides released by enzymatic hydrolysis of hen egg white lysozyme, J Agric Food Chem, 52, 1088–1094. Mine Y, Noutomi T and Haga N (1990), Thermal induced changes in egg white proteins, J Agric Food Chem, 38, 2122–2125. Nakamura R and Ishimaru M (1981), Changes in the shape and surface hydrophobicity of ovalbumin during its transformation to S-ovalbumin, Agric Biol Chem, 45, 2775–2780. Nakamura S, Gohya Y, Losso J N, Nakai S and Kato A (1996), Protective effect of lysozyme-galactomannan or lysozyme-palmitic acid conjugates against Edwardsiella tarda infection in carp, Cyprinus carpio, FEBS Lett, 383, 251–254. Nakamura S, Kato A and Kobayashi K (1991), New antimicrobial characteristics of lysozyme-dextran conjugate, J Agric Food Chem, 39, 647–650. Nakamura S and Kato A (2000), Multi-functional biopolymer prepared by covalent attachment of galactomannan to egg white proteins through naturally occurring Maillard reaction, Nahrung, 44, 201–206. Nakamura S, Ogawa M, Nakai S, Kato A and Kitts D D (1998), Antioxidant activity of a Maillard-type phosvitin–galactomannan conjugate with emulsifying properties and heat stability, J Agric Food Chem, 46, 3958–3963. Nara E, Miyashita K and Ota T (1995), Oxidative stability of PC containing linoleate and docosahexaenoate in an aqueous solution with or without chicken egg albumin, Biosci Biotech Biochem, 59, 2319–2320. Pellegrini A (2003), Antimicrobial peptides from food proteins, Curr Pharm Des, 9, 1225–1238. Pellegrini A, Hülsmeier A J, Hunziker P, and Thomas U (2004), Proteolytic fragments of ovalbumin display antimicrobial activity, Biochim Biophys Acta, 1672, 76–85.
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Fractionation of egg proteins and peptides 617 Pellegrini A, Thomas U, Bramaz N, Klauser S, Hunziker P and von Fellenberg R (1997), Identification and isolation of bactericidal domain in chicken egg white lysozyme, J Appl Microbiol, 82, 372–378. Pellegrini A, Thomas U, von Fellenberg R and Wild P (1992), Bactericidal activity of lysozyme and aprotinin against Gram-negative and Gram-positive bacteria related to their basic character, J Appl Bacteriol, 72, 180–187. Perlman G E (1952), Enzymatic dephosphorylation of ovalbumin and plakalbumin, J Gen Physiol, 25, 711–726. Powrie W D and Nakai S (1985), Characteristics of edible fluids of animal origin: eggs, in Fennema O, Food Chemistry 2nd edition, New York, Marcel Dekker Inc., 829–855. Radomski M W and Cook W H (1964), Fractionation and dissociation of the avian lipovitellins and their interaction with phosvitin, Can J Biochem, 42, 395–406. Roy I, Rao M V S and Gupta M N (2003), An integrated process for purification of lysozyme, ovalbumin and ovomucoid from hen egg white, Appl Biochem Biotechnol, 111, 55–63. Sheumack D D and Burley R W (1988), Separation of lipid-free egg yolk proteins by high-pressure liquid chromatography using solvents containing formic acid, Anal Biochem, 174, 548–551. Sriram V and Ygeeswaran G (1999), Improved recovery of immunoglobulin fraction from egg yolk of chicken immunized with AsialoGM1, Russ J Immunol, 4, 131–140. Stadelman W J and Coterril O J (1973), Egg science and technology, Westport: Avi Publishing, p. 314. Stein P E, Leslie A G, Finch J T, Turnell W J, McLaughlin P J and Carrell R W (1990), Crystal structure of ovalbumin as a model for the reactive centre of serpins, Nature, 347, 99–102. Stein P E, Leslie A G, Finch J T and Carrell R W (1991), Crystal structure of uncleaved ovalbumin at 1.95 Å resolution, J Mol Biol, 221, 941–959. Tsutsui T and Obara T (1984), Preparation and characterization of phosvitin from hen’s egg yolk granule, Agric Biol Chem, 48, 1153–1160. Van Het Schip F D, Salmallo J, Broos J, Ophuis J, Mojet M, Gruber M and Geert A B (1987), Nucleotide sequence of a chicken vitellogenin gene, J Mol Biol, 196, 245–260. Warner R C and Weber I (1951), The preparation of crystalline conalbumin, J Biol Chem, 191, 173–180. Watanabe K, Matsuda T and Sato Y (1981), The secondary structure of ovomucoid and its domain as studied by circular dichroism, Biochim Biophys Acta, 667, 242–250. Watanabe K, Shimoyamada M, Onizuka T, Akiyama H, Niwa M, Ido T and Tsuge Y (2004), Amino acid sequence of a-ovomucin in hen egg white ovomucin deduced from cloned cDNA, DNA Seq, 15, 251–261. Wijers M C, Pouliot Y, Gauthier S F, Pouliot M and Nadeau L (1998), Use of nanofiltration membranes for the desalting of peptide fractions from whey protein enzymatic hydrolysates, Le Lait, 78, 621–632. Williams J, Elleman T C, Kingston I B, Wilkins A G and Kuhn K A (1982), The primary structure of hen ovotransferrin, Eur J Biochem, 122, 297–303. Wiseman R L, Fothergill J E and Fothergill L A (1972), Replacement of asparagines by aspartic acid in hen ovalbumin and a difference in immunochemical reactivity, Biochem J, 127, 775–780. Yamauchi K, Kurizaki J and Sasago K (1976), Polypeptide composition of hen’s egg yolk very low-density lipoprotein, Agric Biol Chem, 40, 1581–1586. Yoshii H, Tachi N, Ohba R, Sakamura O, Takemaya H and Itani T (2001), Antihypertensive effect of ACE inhibitory oligopeptides from chicken egg yolks, Comp Biochem Physiol C Toxicol Pharmacol, 128, 27–33.
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618 Separation, extraction and concentration processes Young L L and Gardner F A (1972), Preparation of egg white ovomucin by gel filtration, J Food Sci, 37, 8–11. Young A C, Tilton R F and Dewan J C (1994), Thermal expansion of hen egg-white lysozyme. Comparison of the 19 Å resolution structures of the tetragonal form of the enzyme at 100 K and 298 K, J Mol Biol, 235, 302–317.
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Supercritical-fluid extraction of lycopene from tomatoes 619
20 Supercritical-fluid extraction of lycopene from tomatoes J. Shi and S. Jun Xue, Agriculture and Agri-Food Canada, Canada, Y. Jiang, The Chinese Academy of Sciences, China and X. Ye, Zhejiang University, China
Abstract: Several process parameters for supercritical CO2 fluid extraction – such as pressure, temperature, flow rates, co-solvent or modifier concentrations, resident time, moisture content, particle sizes, and particle size distribution – have individual or combined effects on the recovery of lycopene from tomatoes. The solubility and bioactivity of lycopene, and the composition of the extract and extraction yield can be affected. Improved processing conditions and reduced cost are required if the extraction of lycopene from tomato materials using the supercritical CO 2 fluid extraction process is to become more economical at low throughputs. Key words: bioactivity, lycopene, supercritical fluid extraction, tomato.
20.1 Introduction Lycopene is a phytochemical responsible for the red pigments found in plants. It is a non-provitamin A carotenoid that plays an important role in the biosynthesis of many carotenoids. Recently, the extraction of lycopene and other carotenoids has attracted attention owing to their biological and physiochemical properties, particularly those which possess natural antioxidant activities. Antioxidants have been associated with disease prevention and reduction (Negre-Salvayre et al., 2006; Papas, 1999; Wu, et al., 1999). Structurally, lycopene is an acyclic, open-chain, C40 polyisoprenoid unsaturated carotenoid having 13 double bonds, of which 11 are conjugated, arranged in a linear array, and it has a molecular formula of C40H56 (Fig. 20.1). Because of this unique molecular structure, lycopene has high antioxidant activity and singlet oxygen quenching ability that is thought to be beneficial
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620 Separation, extraction and concentration processes
Fig. 20.1 Structure of lycopene.
to human health. The antioxidant protective effect of lycopene and tomatoes has been shown in vitro as well as in vivo (Breinholt et al., 2000; Hadley et al., 2003; Rao and Agarwal, 1998; Suganuma and Inakuma, 1999; Willcox et al., 2003). Therefore, foods containing lycopene are of considerable interest. The singlet oxygen quenching ability of lycopene has been found to be three times greater than that of b-carotene and a-tocopherol (Stahl and Sies, 1992, 1996). Epidemiological studies and investigations have shown important roles of lycopene and other carotenoids in free radical inactivation and fat peroxidation inhibition (Krinsky and Rock, 1998). The health benefits of lycopene-rich diets and its effect on minimizing the risk of cardiovascular ailments and various forms of cancers have been reported (Palozza, 1998; Olson, 1986). More recent studies have demonstrated the important functions of lycopene and other carotenoids in the pathophysiology of chronic diseases (Rao and Agarwal, 1999; Rao and Rao, 2007). Increasingly, in vivo and in vitro clinical studies have shown its protective effect against the growth of tumor cells and its ability to protect against cardiovascular (Arab and Steck, 2000) and coronary heart diseases, and cancer (Clinton, 1998). With its 11 conjugated and two non-conjugated double bonds, lycopene is a more efficient antioxidant (singlet oxygen quencher) than b-carotene, a-carotene, and a-tocopherol (Mascio et al., 1989; Shi, 2002; Shi and Le Maguer, 2000). As the most abundant carotenoid in tomatoes, lycopene is the dominant pigment responsible for the color of tomatoes. Lycopene occupies the largest portion, about 80–90%, of the carotenoids in ripe tomatoes, followed by b-carotene, phytoene, and the other minor carotenoids. Concentrations of lycopene in tomatoes approach 50–100 mg kg–1, and it has a higher colour intensity than b-carotene (Shi and Le Maguer, 2000). Other sources of lycopene include pink grapefruit, guava, watermelon, autumn olive, and apricots. Ripe tomato skins contain approximately five times more lycopene than the pulp. The industrial processing of tomato products produces wastes such as seeds and skins. However, reuse of those wastes is limited to animal feed. Even though ripe tomatoes are the most abundant source of lycopene, over 90% is located in the skin, which constitutes the greater part of the waste, and is a potential natural source for lycopene extraction (Shi and Le Maguer, 2000). Lycopene content in plant materials is dependent on the species and the temperature at which growth and maturation occurs. Factors such as climate trends significantly alter the amount of lycopene in the material, thus affecting the yield. Lycopene is the last pigment to appear during maturation and its
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Supercritical-fluid extraction of lycopene from tomatoes 621 development is inhibited by temperatures higher than 30–32 °C, whereas other carotenoids degrade at these temperatures. As a result, the pigment concentrations in plant materials may vary between crop cycles (Shi et al., 1999b). The concentration of lycopene also changes during ripening and storage. Liu et al. (2009) reported that the accumulated lycopene contents of tomatoes did not change significantly during the first four days of storage. However, between days 4 and 21, the lycopene content increased. A growing interest in and demand for healthy, environmentally safe, and cost-efficient products has driven the research and application of new technologies in the food, pharmaceutical and cosmetic industries. Extraction processes are commonly used to enrich and detoxify food through the removal of targeted components from natural products. Organic solvents are used in conventional methods for the extraction of bio-compounds such as lycopene from plant materials. However, these solvents not only generate environmentally hazardous problems and requests for expensive disposal procedures for the chemical extraction solvents, but also chemical residues remaining in the final products become a major safety concern. With increasing government restrictions reflecting consumer concerns on food safety, alternative and reliable extraction techniques are of great interest. Extraction processes are needed for ‘additive-free natural’ products. It is clear that the current concern for safety in food products has increased interest in ‘green’ extraction techniques, instead of the conventional organic solvent extraction processes. The determination of lycopene in food products and development of a safe ‘green’ extraction process to complement the fortification of functional foods is of great public interest. One possible environmentally friendly alternative is supercritical fluid extraction (SFE), in particular using supercritical CO2 ‘green technology’ because it is physiologically harmless, environmentally safe, non-explosive, exhibits high selectivity as a result of low viscosity, high diffusivity, and liquid-like density, as well as being readily available and easily removed from products (Simandi et al., 2002). Because of the CO2 supercritical state (31 °C and 7.38 MPa), procedures should allow supercritical operation of thermally labile compounds that would be easily degraded at high temperatures. SFE has attracted growing interest for the recovery of natural compounds for large-scale industrial production over recent decades (O’Day and Rosenau, 1982). SFE is successfully and widely used for the extraction of lycopene from ripe tomatoes (Cadoni et al., 2000) and tomato processing wastes (Baysal et al., 2000; Kassama et al. 2008; Ollanketo et al., 2001; Rozzi et al., 2002; Sabio et al., 2003; Topal et al., 2006; Vasapollo et al., 2004). One significant thermodynamic advantage of using supercritical fluid is its ease of separation from the extracted solutes by simply modifying the operation conditions, either pressure or temperature. Supercritical fluids have liquidlike densities that give superior mass transfer characteristics compared with organic solvents, and are characterized as low-viscosity and high-diffusivity fluids. In addition, supercritical fluids have low surface tension leading to
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622 Separation, extraction and concentration processes easy penetration into the porous biological matrix of plant material while releasing the targeted components.
20.2 Supercritical-fluid extraction (SFE) of lycopene Supercritical fluids are, by definition, at a temperature and pressure greater than or equal to the critical temperature and pressure of the fluid. They actually have physical properties somewhere between those of a liquid and a gas. Supercritical fluids are able to spread out along a surface more easily than a true liquid because they have lower surface tensions than liquids. The supercritical CO2 is fed at high pressure by means of a pump, which pressurizes the extraction tank and also circulates the supercritical fluid throughout the system. Figure 20.2 demonstrates an example of a typical single-stage supercritical CO2 extraction system. Once the supercritical CO2 reaches the equilibrium state in the extraction vessel, the extraction process proceeds with the manipulation of pressure and temperature to achieve the ideal operating conditions. The mobile phase, consisting of the supercritical CO2 and the solubilized lycopene component, are transferred to the separator where the solvating power of the fluid is decreased by increasing the temperature and or decreasing the pressure of the system. The lycopene extract precipitates in the separator whereas the supercritical CO2 is either released or recycled back to the extractor. As the solution containing the extracts of lycopene compounds leave the extractor and flow to the first separation vessel via the pressure regulator, a paste of lycopene oleoresins settles to the bottom as they separate and this is collected, whereas the remaining solution goes to the second-stage separator where the fractionation of the lycopene components occurs. A third stage of separation might be required for the complete isolation of pure lycopene
Back pressure valve
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Fig. 20.2 Schematic diagram of a typical single-stage supercritical fluid extraction system with CO2 (modified from Shi et al., 2006).
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Supercritical-fluid extraction of lycopene from tomatoes 623 components. In commercial-scale processing, multiple extraction vessels are normally used to enhance operation efficiency and throughput (Fig. 20.3). The process requires an intermittent batch system because it is interrupted at the end of the extraction period, when the pressure must be released so the extraction vessel can be switched out of the extractor loop. By having two or more interchangeable extraction vessels, the unloading and reloading of a vessel can occur while extraction of a previously charged vessel is in progress. This reduces the downtime and improves the overall production efficiency.
20.3 Factors affecting lycopene yield Many studies have successfully extracted lycopene from tomato materials by SFE. The technology has been continuously developed and users have established environmentally sound and complimentary safe techniques for extracting food-grade, bioactive components used as natural ingredients from agricultural materials. There have been numerous studies on the effects of various independent variables – such as pressure, temperature, flow rates, and co-solvent or modifier concentrations – on lycopene yield using SFE (Baysal et al., 2000; Ollanketo et al., 2001; Ciurlia et al., 2009; Kassama et al., 2008; Yi et al., 2009, Shi et al., 2009a; 2009b; Vaughn et al., 2008). Two-stage separation vessels
Storage tank
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Fig. 20.3 Schematic diagram of commercial scale multi-stage supercritical fluid extraction system used to fractionate bioactive components.
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624 Separation, extraction and concentration processes The optimization of pressure, temperature, and flow rate is necessary to obtain a high recovery of lycopene. Baysal et al. (2000) observed the highest recovery (54%) of lycopene obtained at 55 °C, 30 MPa, and 4 kg h–1 with 5% ethanol as modifier. Ollanketo et al. (2001) recovered 94% of lycopene in 15 min at 40 MPa and 110 °C. Rozzi et al. (2002) obtained a maximum recovery (61%) at 80 °C, 35 MPa, and 2.5 mL min–1. Yi et al. (2009) studied the effects of SFE parameters on the recovery of lycopene and antioxidant activity. In this study, the maximum total lycopene content of 31.25 mg g–1 of raw tomato was extracted at the highest temperature of 100 °C, 40 MPa and 1.5 mL min–1. Process parameters such as solvent flow rate, residence time, moisture content, particle sizes, and particle size distribution in conjunction with supercritical pressures and temperatures are key parameters for achieving optimum results. Most of these parameters can have individual or combined effects on the extraction rate, bioactivity, and composition of the extract. For example, the residence time can have an important influence on the composition of the extracted compound. 20.3.1 Effects of temperature, pressure and flow rate on recovery of lycopene Effect of pressure Plate IIa (between pages 292 and 293) shows that pressure significantly influences the rate of extraction. Higher density causes a double effect that increases the solvation power and reduces the interaction between the fluids and matrix, resulting in a decrease in the diffusion coefficient. Excessive pressure also increases the compactness of the sample matrix, thus reducing the pore sizes and apparently reducing the mass transport which eventually diminishes the yield (Tonthubthimthong et al., 2001). It is commonly considered that an increase in pressure results in an increase in CO2 density, increasing the solvating power of the supercritical fluid. Thus, higher pressure is responsible for quantitative recoveries and stronger interactions between the fluid and the matrix (Shi et al., 2009a; Topal et al., 2006). Effect of temperature Temperature is a parameter with significant influence on SFE yield; thus manipulating it could have an adverse implication on the process and yield. Plate IIb shows the general trend of increased extraction yield as temperature increases relative to the pressure. The extraction yields increased both with temperature and pressure, whereas the effect of temperature is more significant than that of pressure. Vági et al. (2007) maximized the recovery of lycopene at 46 MPa and 80 °C by SFE. At 40–100 °C, maximum yields were obtained for extraction of lycopene. The highest temperature gave the maximum lycopene content (Rozzi et al., 2002). Cadoni et al. (2000) reported that at 85 °C a maximum of 65% of lycopene was recovered from the pulp
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Supercritical-fluid extraction of lycopene from tomatoes 625 and skins of ripe tomatoes. However, with further elevation of temperature, degradation and isomerization of lycopene was often encountered (Shi et al., 1999a, 2009b, 2002; Shi and Le Maguer, 2000; Yi et al., 2009). It is well-known that increasing the temperature reduces the solvent density and consequently decreases the recovery of lycopene at constant pressure. At the same time, higher temperature promotes the solubility of lycopene and increases the recovery by increasing mass transfer of lycopene in the matrix, and from the matrix to the CO2 extraction medium (Marsili and Callahan, 1993; Shi et al., 2007d). Therefore, the increase in temperature could have either a positive or a negative effect, as a result of the balance between CO2 density (r) and solute vapour pressure. The increase in the recovery of lycopene by SFE depends more on the solute’s vapour effect. Although lycopene is not stable during long heating times and was approximately 53.5% degraded after 60 min at 100 °C through isomerization or auto-oxidation (Boskovic, 1979; Mayeaux et al., 2006; Shi and Le Maguer, 2000; Shi et al., 1999a). However, in the thermal processing of tomato products, the total lycopene contents still increased in the products owing to the elevated release of lycopene from the tomato tissue matrix (Dewanto et al., 2002; Seybold et al., 2004; Toor and Savage, 2006). For SFE, the enhanced release of lycopene from the insoluble fibre portion of tomato skin may also contribute to the increased recovery. Effect of CO2 flow rate In an extraction, the flow rate of supercritical CO2 has an influence on the extraction efficiency. The flow rate controls the amount of solvent (e.g. CO2) to be used and the extraction time. Extraction is a diffusion-based process, with the solvent required to diffuse into the matrix, and the extracted component to diffuse out of the matrix into the solvent at equilibrium state. The flow rate also greatly affects solubility of extracted components in fluid and equilibrium of diffusion between solvent and extracted components during extraction. If the concentration of extracted components is much higher than the solvent, the solubility of extracted components in the solvent could be limited and slow down the extraction process. Therefore, optimization of flow rate is important for extraction time and cost efficiency. For example, to maximize the rate of extraction, the flow rate should be high enough for the extraction to be completely diffusion limited, but this is very wasteful of solvent. In contrast, the lower flow rate would minimize the amount of solvent used, but the extraction should be completely solubility limited, and thus would take a very long time. Therefore, the optimum flow rate is probably somewhere in the region where both solubility and diffusion are significant factors. At optimum temperature and pressure, the extraction process is performed under conditions in which the CO2 flow rate is variable. The linear correlation is observed for the flow rates and the extraction rates. The extraction rate increased with increasing flow rate. When the intra-particle diffusion
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626 Separation, extraction and concentration processes resistance is considered as a factor in the material matrix, more CO2 is required to make contact with the targeted components. Thus, the extraction rate, is a function of CO2 consumed: the greater the amount of supercritical gas used, the higher the yield obtained (Friedrich and Pryde, 1984). When the temperature, pressure, and overall extraction time are kept constant, the effect of various solvent flow rates on the carotenoid yield is minor. Plate IIc shows the effect of flow rate on total lycopene yield at constant pressure (30 MPa) or temperature (70 °C). An increase in flow rate from 1.0 to 2.0 mL min–1 did not show a significant change in the yield of lycopene (Yi et al., 2009). Results indicated that the higher flow rate could produce faster extractions and higher recoveries since the higher flow rate could overcome the interface resistance for lycopene transport from the tomato skin tissue to the CO2 fluid (Hawthorne et al., 1995). However, excessively high flow rates would produce undesirable results owing to a decrease in contact times between the solvent and the solute (Topal et al., 2006). This may force the solvent to leave the system without reaching its solubility limit (Anitescu et al., 1997). High flow rates may also cause the sample to compact and restrict the fluid flow, thus reducing the amount of solvent that comes in contact with the solute (Franca and Meireles, 2000). Increasing the CO2 flow rate would increase fluid diffusivity and improve the heat and mass transfer within the system, because the resulting fluid temperature approaches the wall temperature more quickly (Shi et al., 2007a). The effect of solvent flow rate on mass transfer rate and total yield in the constant extraction rate period is significant. Effect of operating time Residence time is an important factor that influences yield and the economic viability of the process. When optimizing extraction conditions, the extraction time of the process is an important parameter. For the extraction of lycopene by SFE Baysal et al. (2000) found that the highest carotene yield was obtained at an extraction time of 2 h, as opposed to 1 or 3 h. A 1-h extraction may not suffice for the maximum amount of carotenoids to be dissolved in the solvent, whereas over 3 h there is an increase in the occurrence of degradation. The duration and intensity of thermal processing are directly correlated to the degree of isomerization and degradation of the targeted compounds (Baysal et al., 2000). Therefore, to minimize the degradation effect, it is advisable to reduce the extraction time as much as possible. In addition, optimizing extraction conditions towards the shortest extraction time possible with the maximum extraction recovery is ideal from the point of view of the extract quality and processing cost. Effect of moisture content of raw materials Water represents approximately 80–90% of the total weight of fresh plant materials and interferes with the effectiveness of the SFE (Hopper and King, 1991). Thus, the initial moisture content may improve or impede the
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Supercritical-fluid extraction of lycopene from tomatoes 627 yield results depending on the polarity of the compound. The non-polar compounds such as lycopene produce smaller yields with higher moisture content (Anitescu et al, 1997; Chao et al., 1993). Water has a small but finite solubility in supercritical CO2, therefore, it can also be extracted with the targeted components. The presence of a little water (less than 10%) aided in the extraction of lycopene, particularly at the beginning of the process (Shi et al., 2006). A moisture content of 7–18% had negligible effect on extractability of oil. However, above an 18% moisture content, the presence of water affected the extraction and the efficiency decreased (Nagy and Simándi, 2008). The high water content in the matrix inhibited the flow of SCF by changing surface tension and contact angles as a result of phase interaction. Therefore, water removal in most cases frees the internal pores and thus increases the mass transport intensity. Higher moisture contents cause a higher probability of formation of a thin film of water between the matrix and the SCF phase. Possible effects of water include an increase in distance that carotenoids must travel to reach the solvent and swelling of the cell matrix, which have negative and positive effects on the process, respectively. Effect of particle size Particle size is a factor that influences the extraction recovery, having a significant impact on the flow behaviour of SCF in the sample matrix. The extraction process and diffusivity of the extraction medium are hindered by large particle sizes (>550 mm) because of the distance which the lycopene travels from the inside of the particle to the surface (Shi et al., 2007a). By reducing the size of the particles, components such as carotenoids have less distance to travel. Thus, the collective amount of surface area with which the solvent is in contact is increased, an important factor during the initial stages of extraction (Ollanketo et al., 2001). The particle size is affected by sample pre-treatment. For example, the methods of drying (air, oven, vacuum, or freeze drying), milling, and other mechanical, ultrasonic, high-pulsed electronic field, and non-mechanical treatments can change the particle size by attrition or size reduction. Numerous experimental results have indicated that smaller particle sizes are ideal for extraction processes, and the manner in which the solid material is crushed and powdered also affect the results (Ferreira and Meireles, 2002; Spanos et al., 1993). The degree to which the material is crushed determines the amount of cells that are broken, and this has a positive correlation with recovery. The smaller the particle size, the larger is the surface area, and the bioactive components are released more easily. However, if the particle size is exceedingly small that can favour channelling which commonly results in a decrease in recovery (Ziémons et al., 2005). Only 40% lycopene recovery was obtained from powdered tomato skins crushed with a household blender, relative to the amount achieved when sea sand was used to assist the grinding procedure (Ollanketo et al., 2001). Therefore, optimizing particle
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628 Separation, extraction and concentration processes size significantly improves the yield of extraction. Sun and Temelli (2006) noted that the optimal range of particle size was 0.25–0.5 mm for lutein, b-carotene, and a-carotene. However, for certain processes, it is important to be cautious of the degree of crushing. Gouveia et al. (2006) studied the effects of slight, moderate and complete crushing, which gave 20%, 40%, and 65% carotenoid recovery, respectively, and also found that samples that have been extremely crushed may not be affected by the addition of entrainers.
20.4 Effects of pressure and temperature on the antioxidant activity of lycopene Recovery is increased with increasing temperature owing to increasing solubility of the targeted compounds; therefore, for economic reasons, food industries prefer to increase the temperature to achieve higher yields. However, the bioactivity of the lycopene may be compromised by heat because of the instability of lycopene at higher temperatures (Mayer-Miebach et al., 2005; Shi et al., 2002b, 2004b, 2008, 2009b; Vega et al., 1996; Yi et al., 2009). Yi et al. (2009) reported that the recovery and bioactivity of lycopene increased when the temperature increased from 20 to 60 °C. The recovery of lycopene increased slowly and its antioxidant activity declined above 80 °C. Pol et al. (2004) made a remarkable observation that lycopene recovery was increased with increasing temperature, but the bioactivity of the extracted lycopene did not follow the trend of the recovery of lycopene extracts at high temperature. No significant effects of pressure or flow rate on the antioxidant activity were observed (Yi et al., 2009). Cadoni et al. (2000) and Shi and Zhou (2006) reported that the results obtained from SFE of lycopene from tomatoes are poorly reproducible because of the decomposition and isomerization of lycopene during the extraction process. Yi et al. (2009) investigated the effects of the two major supercritical fluid extraction parameters (temperature and pressure) on the antioxidant activities of the extracts as shown in Fig. 20.4. The recovery increased with the increase in pressure, whereas antioxidant activity decreased slightly. The recovery of total lycopene increased with temperature, but the antioxidant activity of the lycopene-rich extract decreased with increasing temperature (Hackett et al., 2004; Yi et al., 2009). At 25 and 50 °C, lycopene in an oleoresin is degraded predominately through oxidation, whereas isomerization increases at 75 and 100 °C. If lycopene is isomerized, the proportion of lycopene isomers is changed, that then further influences the antioxidant activity. In general, the change in antioxidant activity in lycopene-rich extracts is mainly influenced by their lycopene content, proportions of the isomers of lycopene, and the synergistic effects of carotenoids in the extract (Shi et al., 2004a,b, 2004c, 2007c; Shixian et al., 2005). The total antioxidant activity of
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0.20
20 0.15
15
0.10
10 5 20
25
30 Pressure (MPa) (b)
35
Torlox of lycopene (mM mg–1)
35
Torlox of lycopene (mM mg–1)
Lycopene concentration (mg g–1)
Supercritical-fluid extraction of lycopene from tomatoes 629
0.05 40
Fig. 20.4 Changes of lycopene yield (mg –1g raw material) (d) and antioxidant activity (mM mg–1 Torlox) of the extract (m) (a) at 40 to 100 °C, 30 MPa and flow rate of 1.5 mL min–1, and (b) at 20–40 MPa, 70 °C and flow rate of 1.5 mL min–1 (modified from Yi et al., 2009).
the extract, that is the antioxidant activity of each unit of lycopene multiplied by the total lycopene amount in the extract, firstly significantly increases as the temperature increases, and then slightly decreases as temperature is kept at a high temperature. The antioxidant activity of extracts first increased with increasing lycopene contents in the extracts between 40 and 80 °C, and then started to decline thereafter. The slight decrease above 90 °C might be affected by the degradation of other carotenoids (Cocero et al., 2003). Dewanto et al. (2002) found that thermal processing increased the total antioxidant activity of tomato from 4.13±0.36 to 6.70±0.25 mmol g–1 of vitamin C equivalents after 30 min of heating at 88 °C. The total antioxidant activity probably increases because larger amounts of lycopene are released from the matrix during thermal processing. HPLC chromatograms of the extracts
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630 Separation, extraction and concentration processes obtained at 50 and 90 °C are shown in Fig. 20.5. All-trans lycopene is the major configuration eluting with a retention time of 15.87 min under both conditions. Three major cis-isomer peaks are detected around 6.8, 8.9, and 16.1 min of elution time in extracts prepared at 50 °C. When the extraction temperature is elevated to 90 °C, total lycopene increased 1.18 fold, whereas all-trans lycopene contents increased only 1.03 fold and total cis-isomers increased more than 2.13 fold, respectively. Moreover, more than three major peaks of cis-isomers are detected (Fig. 20.5) because of new cis-isomer generations. The higher increase in cis-isomers than total and all-trans lycopene indicates the transformation of all-trans to cis-isomers. As extraction vessel temperature increased from 50 to 90 °C, the ratio of all-trans to total cislycopene isomers changed from 1.67 to 1.40 (Table 20.1). Further increases in the operating temperature resulted in dramatic isomerization. The lower ratio of trans- to cis-isomers obtained at high temperature indicated higher lycopene bioavailability owing to the cis-isomers generated (Boileau et al., 1999). The isomerization of lycopene in tomato samples during thermal processing above 75 °C, and the decrease in the trans- to cis-isomer ratios from 9:1 to 2:3 after heating at 140 °C, were demonstrated (Cocero et al.,
4 cis isomers mAU
3
trans-lycopene
2 5-cis
1 0
0
5
10
15 20 Time (min) (a)
25
30
25
30
trans-lycopene
4 mAU
cis isomers 3 2
5-cis
1 0
0
5
10
15 20 Time (min) (b)
Fig. 20.5 Chromatograph of SFE extract obtained at (a) 50 °C and (b) 90 °C at 30 MPa and flow rate of 1.5 mL min–1 (modified from Yi et al., 2009). mAu, milli absorbance unit. © Woodhead Publishing Limited, 2010
Supercritical-fluid extraction of lycopene from tomatoes 631 Table 20.1 All-trans and cis-isomers contents and ratio in the SFE extract obtained at 40–100 °C, 30 MPa and a flow rate of 1.5 mL min–1 (modified from Yi et al., 2009) Temperature Total content All-trans (mg g–1) content (mg g–1) 100 90 80 70 60 50 40
29.46 26.12 22.69 14.66 13.85 12.00 10.35
16.75 15.26 13.76 9.04 8.74 7.50 6.52
Total cis Ratio Ratio Ratio* content (trans:total) (cis:total) (trans:cis) (mg g–1) 12.71 10.86 8.93 5.62 5.11 4.50 3.83
0.57 0.58 0.61 0.62 0.63 0.63 0.63
0.43 0.42 0.39 0.38 0.37 0.37 0.37
1.32c 1.40c 1.54bc 1.62ab 1.71a 1.67a 1.70a
*The different letters shows the values that are significantly different (P <0.05) (n = 3).
2003; Mayer-Miebach et al., 2005; Shi et al., 2002a; 2002b; Xianquan et al., 2005).
20.5 Effect of co-solvent and modifiers in lycopene extraction The use of co-solvents or modifiers during SFE is key to enhancing the extraction efficiency and cost effectiveness of the processes. Joslin et al. (1996) indicated two significant attributes of co-solvents: the interaction between the co-solvents or modifiers with the solute (direct effect) and the co-solvents or modifiers with solvent interactions (indirect effect). Co-solvents or modifiers when used in small doses (1 to 5% mol) can change the overall characteristics of the extraction fluid in terms of polarity, solvent strength, and specific interactions. These changes, in turn, can significantly alter the density and compressibility of the SCF. Additionally, they can improve selectivity for desired components and facilitate selective fractional separations. The effect of a modifier depends upon the nature of the solute to be extracted (Walsh et al., 1987). The first basis for co-solvent or modifier selection is the increased solubility of the target in the modified CO2 fluid (Pourmortazavi and Hajimirsadeghi, 2007; Shi et al., 2007b, 2009a). Carbon dioxide has adequate solvent properties for extraction of targeted compounds. The addition of a small amount of a liquid modifier or co-solvent such as water, oil, or ethanol can significantly enhance the extraction of non-polar compounds such as carotenoids. Another advantage of selecting a co-solvent or modifier is the ability to distort and swell the matrix, favouring the penetration of the CO2 into the matrix in order to extract the analyte (Casas et al., 2007). Ethanol, water, and vegetable oil under various temperatures and pressures are often used. The addition of these co-solvents or modifiers significantly increases the total recovery because the solvent power of supercritical CO2 could be © Woodhead Publishing Limited, 2010
632 Separation, extraction and concentration processes increased by the addition of small amounts of co-solvents or modifiers. The effect of co-solvents or modifiers is dependent on their concentration in the supercritical phase, and this is determined by the phase behaviour of the mixture under operating conditions. It also should be noted that the addition of large amounts of modifier will change the critical parameters of the mixture (Pourmortazavi and Hajimirsadeghi, 2007). 20.5.1 Effect of ethanol as modifier Baysal et al. (2000) used ethanol at different concentrations (5, 10, and 15% w/w) to recover b-carotene and lycopene from tomato paste. The addition of ethanol increases the bulk density of supercritical CO 2 because of the higher density of the co-solvent and clustering of supercritical CO 2 molecules around the co-solvent (Güçlü-Üstündag and Temeli, 2005). The solubility of trans-b-carotene both in pure and ethanol-modified supercritical CO2 increases with temperature from 40 to 60 °C (Sovová et al., 2001). The ethanol modifier increased the solubility by one order of magnitude and the increase in solubility was proportional to the square root of the modifier concentration (Güçlü-Üstündag and Temeli, 2005; Shi et al., 2009a, 2009b; Sovová et al., 2001). Because of the high molar mass and elongated shape of the lycopene molecule, the modifier dilutes the extract, reducing the viscosity, and thereby enhancing the flow of the extract through the extractor. The increase in recovery showed a linear correlation with the increase in ethanol concentration at 45 °C. However, the increased recovery (6%) was lower when the ethanol concentration was increased from 10 to 15% than the increase (11%) when the ethanol concentration was increased from 5 to 10% (Shi et al., 2009b). The results suggest that selecting the proper concentration of modifier could improve extraction efficiency and reduce the operating cost. 20.5.2 Effect of water as a modifier Water enhances the analyte–modifier–matrix interaction because water opens the pores and causes the matrix to swell thereby allowing the supercritical fluid better access to the analytes to bring them out of the matrix. The recovery of lycopene increased significantly with an increase in water concentration in tomato skin materials (Yi et al., 2009). An excess of water (greater than 18%) causes mechanical difficulties such as the extraction fluid (e.g. CO2) clogging the restrictor (Casas et al., 2007) because the high water content makes material transport difficult. Previous studies reported that the extraction of lycopene from tomatoes varied in terms of the corresponding moisture content. Baysal et al. (2000) extracted lycopene from tomato paste waste having 24% moisture content, whereas Rozzi et al. (2002) extracted lycopene from tomato waste materials with a moisture content of 48.4%. As the moisture content of the tomato skins in the reported experiment was
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Supercritical-fluid extraction of lycopene from tomatoes 633 7%, the addition of 15% water would be approximately equal to a moisture content of 19% (Yi et al., 2009). Even though water is only about 0.3% soluble in supercritical CO2 (Lehotay, 1997), water has been used as a cosolvent in a number of SFE applications. In addition, water aids the extraction process by increasing the polarity of the supercritical fluids, thus enabling higher recoveries of relatively polar species (Pourmortazavi and Hajimirsadeghi, 2007). The adsorption of water onto the polar sites of the plant material matrix weakens the bond between the extractable component and the matrix, leading to an increased vapour pressure of the extract and enhanced rate of desorption (Brady et al., 1987). 20.5.3 Effect of edible oil as a modifier Considering the lipophilic properties of lycopene, edible oils used as alternative modifiers were proposed to enhance SFE (Krichnavaruk et al., 2008; Sun and Temelli, 2006; Vasapollo et al., 2004). One advantage is that such edible oil does not need to be subsequently separated from the product. Examples of the use of vegetable oil as a co-solvent for SFE include the extraction of lycopene from tomatoes using hazelnut oil (Vasapollo et al., 2004); the extraction of carotenoids from carrots using canola oil (Sun and Temelli, 2006), and the extraction of astaxanthin from Haematococcus pluvialis using vegetable oil (Krichnavaruk et al., 2008). Vasapollo et al. (2004) tested the effect of 10% hazelnut oil on extractable lycopene content from dried tomato (6% of moisture, average particle size of about 1 mm) under varying pressures and at two different temperatures (40 and 60 °C). Recovery of lycopene with olive oil as a modifier was higher than those obtained with ethanol as a modifier (Shi et al., 2009b). Olive oil, and most probably other vegetable oils, are good solvents for lycopene in the liquid state as they enhance the extraction yield by increasing the solubility of lycopene in the modified supercritical fluid. The slight reduction in recovery obtained with 15% olive oil as a modifier may be attributed to an excess of the oil plugging or clogging the pores of the support. 20.5.4 Effect of binary and ternary modifiers Very little information is available on the influence of the behaviour of two combined modifiers (ethanol and water, water and oil, and ethanol and oil) on the extraction of lycopene. Shi et al. (2009b) compared the effects of binary and ternary modifiers such as water, ethanol, and olive oil on the recovery of lycopene by SFE under various conditions. The results showed that the addition of various quantities of a mixture of ethanol and water resulted in higher lycopene contents than for the single modifier in equal amounts. The addition of a mixture of ethanol and olive oil showed a significant synergistic effect on the recovery compared with the addition of a mixture of ethanol and water, and the enhancement increased with concentration of both ethanol and
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634 Separation, extraction and concentration processes olive oil (Table 20.2). The highest recovery 73.3% was achieved by adding 10% ethanol and 10% olive oil. Olive oil has H-bonding ability, which could link to ethanol and contribute to the interaction between modified supercritical CO2 and the analyte. Owing to the fat-soluble characteristic, the addition of olive oil enhanced the solubility in supercritical CO2 modified with olive oil. Moreover, the addition of ethanol could remove more olive oil and lycopene. The recoveries obtained with a modifier mixture of water and oil were higher than that obtained by modifier mixture of water and ethanol, and lower than that obtained by a modifier mixture of oil and ethanol. The combined water and oil mixtures protect the lycopene extract from oxidative degradation (Boon et al., 2008). 20.5.5 Effects of modifiers on the ratio of all-trans to cis-isomers in lycopene extract According to Mayer-Miebach et al. (2005), there is not a dramatic effect on the isomerization of lycopene below 75 °C. In contrast, all-trans-lycopene was degraded by up to 25% when synthetic lycopene-containing water-in-oil emulsions were heated at temperatures even below 70 °C (Mayer-Miebach and Spiess, 2003). The proportions of cis-isomers in all the extracts obtained with olive oil as a modifier increased compared with extracts obtained with modifier of ethanol or water because the cis-isomers are less likely to crystallize and are more efficiently dissolved in lipophilic solutions such as oil (Shi et al., 2009b).
20.6 Solubility of lycopene in supercritical fluids Solubility refers to the amount of solute that will dissolve in a given amount of solution at thermodynamic equilibrium. Interactions between solvent and Table 20.2 Effects of modifiers on the ratios of all-trans to cis-lycopene in SFE from tomato skins at 350 MPa and 45 and 75 °C (E: ethanol; W: water; O: olive oil) (modified from Shi et al., 2009b) 45 °C
15%E 15% W 15% O 10% E + 10% W 10% E + 10% O 10% W + 10% O 5% E + 5% W + 5% O
75 °C
Recoveries (%)
Ratio (all-trans:cis)
Recoveries (%)
Ratio (all-trans:cis)
42.9 33.3 38.7 39.4 47.0 36.6 37.1
38.2 29:1 33:2 34:2 41:2 31:2 31:2
51.7 48.8 58.2 62.5 73.3 56.8 57.9
45:2 42:2 49:3 54:3 61:3 49:3 49:3
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Supercritical-fluid extraction of lycopene from tomatoes 635 solute molecules influence the tendency of a particular substance to dissolve in a solvent (Shi et al., 2007a). The solubility of targeted compounds is important when establishing the extraction parameters. To obtain the maximum recovery from a source material, the solubility of a component in the solvent should be as high as possible. It also dictates the amount of solvent necessary to optimize the extraction efficiency. The insolubility of lycopene in water causes a problem when attempting to extract it from natural materials. Although it is soluble in organic solvents such as benzene, chloroform and methylene chloride, these are toxic and therefore not ideal candidates for pharmaceutical, food, or cosmetic purposes. This is one of the main reasons why lycopene solubility in supercritical fluids has become of great interest (Cadoni et al., 2000). 20.6.1 Factors affecting the solubility of lycopene in SFE There are numerous factors which affect the solubility of lycopene in supercritical fluids. The solvating power of fluids in the supercritical phase is very sensitive to changes in temperature and pressure. Other parameters, including the presence of co-solvents and modifiers and compound morphology, also have to be considered. Pressure An increase in pressure at a constant temperature increases the solvating power of the fluid. This allows extraction of a wider variety of carotenoids, and thus results in an increase in solvent density, which can lead to a rapid increase in the solubility of the targeted component. Because density increases with pressure, solubility tends to increase with pressure. If the solvating power of the supercritical fluid is increased, it may contribute to a decrease in the diffusion coefficient at higher density, and the interaction between the fluid and the matrix may be reduced (Gordon and Bauernfeind, 1982). Johannsen and Brunner (1997) demonstrated that the solvating power of CO2 increases as the density increases, and found that the increase in solubility of b-carotene in supercritical CO2 occurs when the pressure of the system is increased. Corresponding to this trend, the amount of required CO2 is reduced as the extraction pressure increases. Sabio et al. (2003) noted that, after a certain point, the change in recovery with increasing pressure/density was insignificant. Subra et al. (1997) indicate that at lower pressures, the uncertainty of solubility measurements increases. However, low solubility of the condensed phase, particularly those of b-carotene, occurs at pressures close to the critical pressure of the solvent (Sakaki, 1992). The effect of pressure is enhanced by the addition of modifiers such as canola oil. The solubility of carotenoids increases as the pressure increases; however, modifiers enhance this effect under the same experimental conditions (Shi et al., 2009b; Sun and Temelli, 2006). A decrease in solubility from an increase in pressure at a constant
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636 Separation, extraction and concentration processes temperature is a phenomenon that has been observed in other experiments (Calvey and Page, 1990; Luque de Castro et al., 1994; Tsai et al., 2007). The density of CO2 rises when the pressure is increased; therefore, this may cause a reduction in the diffusivity of the solute through the solvent. Tonthubthimthong et al. (2001) suggested that an increase in pressure may cause the plant material to compress or to become more compact, thus restricting the solvent’s ability to enter the matrix. Pressure may affect the selectivity of the solvent for particular solutes, meaning that a particular range of pressures may favour the extraction of lycopene whereas other conditions are more suitable for other compounds. Figure 20.6 shows the relationship between the solubility of lycopene in supercritical CO2 and pressure at constant temperatures. This relationship suggests that the optimum solubility conditions are not at the highest combination of temperature and pressure tested. However, there is a significant relationship between temperature and lycopene solubility. Increases in temperature result in higher solvent volatility, and the solubility of lycopene is increased as a result. Although there were decreases in solvation power when the pressure was increased at a relatively high temperature, the magnitude of the change in solubility became greater with increasing temperature. In the study of the effects of temperature and pressure on the solubility of lycopene in supercritical CO2 fluid by Shi et al. (2009a), the results showed a considerable decrease in solubility from 25 to 30 MPa at 75 °C. However, the solubility is greater than that at 30 MPa and 65 °C. This may be because the density of the solvent is reduced, making it easier for the lycopene to diffuse into the solid. In addition, the magnitude of the solvent density changes is smaller at elevated pressures. The drop in
Solubility, y2 ¥ 106 (mol solute/106 mol mixture)
2.5
2
1.5
1
0.5
0 100
200
300 Pressure (bar)
400
500
Fig. 20.6 Solubility isotherms for lycopene: (®) 50 °C; () 60 °C; () 70 °C; (d) 80 °C. Lines represent trends and symbols represent experimental data points (modified from Shi et al., 2009a).
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Supercritical-fluid extraction of lycopene from tomatoes 637 solubility from 25 to 30 MPa at 65 and 75 °C may occur because the change in density of supercritical CO2 is not enough to compensate for any matrix packing of the source material that occurs (Shi et al. 2009a). The degree of the change in CO2 density as the temperature increases may be influenced by the pressure. However, the solubility at 75 °C is higher than that at 65 ºC at 30 MPa, despite a relatively severe drop in solubility when the pressure was raised, probably because the increased molecular interactions between the supercritical CO2 and lycopene eclipses the inhibiting effect caused by the increased solvent density. Temperature Elevations in temperature instigate changes in the density of supercritical CO2 and lycopene vapour pressure, and those changes can affect the solubility of the lycopene in a supercritical fluid. The relative effects of solvent density changes and solute vapour pressure may determine the degree that lycopene solubility is enhanced or impeded. The solute vapour pressure has an impact on the solubility because a crystalline solid has a lower solubility than an amorphous solid owing to the difference in free energy, and a relatively higher enthalpy of fusion (Hansen et al., 2001). An increase in temperature raises the solvent’s vapour pressure and decreases the SCF density, but the magnitude of the change in density becomes smaller at elevated temperature (Marentis, 1998). This behaviour may be explained by a complex balancing effect between the density of the solvent and the solute vapour pressure as the temperature is increased (Spanos et al., 1993). At a constant density, solubility increases with a rise in temperature, owing to the increase in vapour pressure of the solid (Johannsen and Brunner, 1997). Rozzi et al. (2002) and Hansen et al. (2001) suggest that the uncertainty of solubility measurements increases as the temperature decreases. The solubility of lycopene increased as the temperature rose at a constant pressure. The degree of increase in solubility of lycopene extracts varied over different temperature ranges. For example, the increase in solubility when the operating temperature rose from 45 to 55 °C, was less than the solubility increase from 55 to 65 °C. Moreover, the increase in solubility of lycopene extracts in CO2 fluid when the temperature increased from 55 to 65 °C was higher than the increase in solubility when from 65 to 75 °C (Shi et al., 2009a). As shown in Fig. 20.6, the solubility of lycopene from 45 to 55 °C increased along with pressure. A crossover point was noted and the magnitude of solubility increase at 30 MPa with temperature was lower than that at 20 and 25 MPa. Topal et al. (2006) extracted lycopene from dried tomato skins using supercritical CO2 and observed a similar trend. Co-solvents and modifiers Extraction can be enhanced using co-solvent or modifier that is able to interact with the target compounds. The addition of an appropriate cosolvent or modifier also influences the solubility of target compounds in
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638 Separation, extraction and concentration processes the supercritical fluid. Although the nature of CO2 drastically reduces or eliminates the solubility of many polar molecules, this is easily remedied by the addition of polar co-solvents or modifiers. Modifiers, such as methanol, increase the solubility of polar compounds but may also increase the critical temperature of the solvent. This may be a problem for compounds that are thermally labile, thus careful selection of supercritical CO2 incorporating an appropriate co-solvent or modifier is required. Another possible drawback of addition of modifiers is product contamination. This may be solved by the addition of a separation step subsequent to the extraction process (Hansen et al., 2001). Chandra and Nair (1997) indicated that the addition of co-solvents or modifiers to supercritical CO2 increase the solvating power. Chang and Randolph (1989) observed that the solubility of b-carotene increased with the addition of toluene. In their study, the results showed that the highest solubility was more than nine times higher than the extraction without toluene under the same conditions, thus showing a large b-carotene solubility enhancement. The influence of various co-solvents and modifiers at a constant temperature, pressure, and solvent flow rate was studied by Ollanketo et al. (2001). Ethanol is the most viable modifier that may be used in largescale lycopene extraction and other food applications. Cygnarowicz et al., (1990) also reported that ethanol functioned as a more effective modifier than methanol or methylene chloride. Modifiers that are miscible in water give higher recoveries. However, a disadvantage of using ethanol as a cosolvent is that an additional step must be added to the process in order to remove it from the final product. This procedure requires the use of heat, thus posing the risk of isomerization of lycopene and other carotenoids (Sun and Temelli, 2006). Natural co-solvents or modifiers used in the extraction of carotenoids include vegetable oils. Sovová et al. (2001) indicated that the modifier effect of vegetable oil was not as great as ethanol, with the increase in b-carotene solubility being four times smaller. The effect of vegetable oil as a modifier is limited owing to its low solubility in CO2. Unlike vegetable oil, canola oil containing a high degree of unsaturated fatty acids compared with the vegetable oil, has been suggested as a natural co-solvent, or modifier, and it is much more soluble in supercritical CO2 than b-carotene is; results show that the carotenoid concentration nearly doubled with canola oil addition. In addition, the effect of increasing the pressure was heightened by this modifier. Sun and Temelli (2006) attributed this effect to interactions between the modifier and the solid material. Penetration of the oil into the plant material matrix may stretch the cell walls, thus making it easier for the supercritical CO2 to penetrate the matrix. In addition, the oil may help release carotenoids in the matrix by loosening the cell structures and increasing solute exposure.
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Supercritical-fluid extraction of lycopene from tomatoes 639
20.7 Conclusion and future trends One of the most important trends in the food industry is the demand for all-natural food ingredients that are free of chemical additives. Natural food antioxidants are derivatives of plant by-products. SFE with CO2 is a viable alternative to the conventional solvent extraction technique to extract bioactive components from agricultural materials. Supercritical CO2 is nontoxic, safe, inexpensive, and has a lower critical temperature (31 °C) that can be easily reached. Therefore, SFE leaves a safe and non-toxic residue in the extracts whereas conventional methods use organic solvents and high temperatures. The mild conditions that accompany with SFE allow the extraction of lycopene and other thermally labile bioactive components from natural sources with minimal damage to the integrity or stability of lycopene and other carotenoids. Large-scale SFE has become a practicable process for the extraction of lycopene from tomato materials. The solvating power of supercritical fluids is sensitive to temperature and pressure changes, thus the extraction parameters may be optimized to provide the highest possible recoveries of lycopene with maximum antioxidant activity. Low viscosity and high mass-transfer rates are additional advantages of using SFE for food products. Ethanol and edible oil as co-solvents or modifiers have been incorporated in lycopene extraction to enhance extraction efficiency and are viable for use in the food industry. SFE offers a unique advantage of adding value to agricultural waste by extracting the lycopene from tomato skins and using it for the fortification of foods and other applications. Its drawbacks are the difficulties in extracting polar compounds and its susceptibility to extracting compounds from a complex matrix where the phase interaction with the intrinsic properties of the product inhibits its effectiveness. Some drawbacks can be ameliorated by using co-solvents. However, much additional investigation is required to understand the solvation effects on targeted bioactive components being extracted. There are many factors which influence the solubility of lycopene and extraction yield in SFE. The fluid (e.g. CO2) density, operating pressure and temperature, as well as flow rate of fluid affect the quality of the extracts and the efficiency of the extraction. By understanding the effect of the parameters that influence extraction, conditions can be chosen to optimize the recovery and cost efficiency. When determining the parameters that should be used to maximize the recovery of lycopene and the solubility, many studies attempted to optimize the operating conditions with incorporated modifier or co-solvent that may be applicable in large-scale applications. For example, non-toxic co-solvents and modifiers that would be acceptable for food processing. Therefore, canola oil and ethanol was used to enhance the extraction yields. The nature of the material used as a source of lycopene governs the availability of the compounds for the extraction process. Although high temperatures in the extraction process generally increase the solubilities of components in supercritical fluids, the conditions under
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640 Separation, extraction and concentration processes which thermally labile targeted compounds are negatively affected should be considered. The intensity and the length of heat processing affect the degree of lycopene isomerization, thus ideally the extraction time and temperature should be minimized. Minimizing such conditions also leads to a more economically viable process. Excessively high flow rates may reduce the contact time between the target components and the solvent and restrict the fluid flow in the samples if it becomes compacted. The optimal flow rate appears to vary with the targeted molecule, the relatively high flow rates having a negative effect on some components. Therefore, to successfully implement SFE technology on an industrial scale, it is necessary to understand the technology, focusing on the proper design of plant components and optimization of the process parameters that provide maximum recovery and quality with minimize operating costs. SFE technology is available in the form of a single-stage batch process, which could be upgraded to a multistage semi-continuous batch, coupled with a multi-separation process to reduce extraction time and increase extraction efficiency. The need to improve the design into continuous modes is growing. Although the capital cost of SFE plant is higher than the traditional or conventional organic solvent extraction plant, the operating costs are lower. However, for a proper comparison of the capital cost of SFE plant versus conventional extraction plant, it is necessary to take into account all the associated equipment used in the conventional extraction processes, such as distillation or evaporation for removal of solvent from the final extracts. On the other hand, the selling price of the products obtained by SFE are also higher because the extracts are safe and natural without chemical contamination. Further development of SFE is necessary for largescale production to be cost effective. Therefore, with improved processing conditions and reduced cost, lycopene extraction from tomato materials by SFE should become even more economical at low throughputs.
20.8 References Anitescu, G., Doneanu, C. and Radulescu, V. 1997. Isolation of coriander oil: comparison between steam distillation and supercritical CO2 extraction, Flavour and Fragrance Journal, 12, 173–176. Arab, L. and Steck, S. 2000. Lycopene and cardiovascular disease. American Journal of Clinical Nutrition, 71, 1691–1695. Baysal, T., Ersus, S. and Starmans, J. D. A. 2000. Supercritical CO2 extraction of b-carotene and lycopene from tomato paste waste. Journal of Agriculture and Food Chemistry, 48, 5507–5511. Boileau, A. C., Merchen, N. R., Wasson, K., Atkinson, C. A. and Erdman, J. W. 1999. Cis-lycopene is more bioavailable than trans-lycopene in vitro and in vivo in lymphcannulated ferrets. Journal of Nutrition, 129, 1176–1181. Boon, C. S., Xu, Z., Yue, X., McClements, D. J., Weiss, J. and Decker, E. A. 2008. Factors affecting lycopene oxidation in oil-in-water emulsions. Journal of Agricultural and Food Chemistry, 56(4), 1408–1414.
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Supercritical-fluid extraction of lycopene from tomatoes 641 Boskovic, M. A. 1979. Fate of lycopene in dehydrated tomato products: carotenoid isomerization in food system. Journal of Food Science, 44, 84–86. Brady, B. O., Chien-Ping, C. K., Dooley, K. M., Knopf, F. C. and Gambrell, R. P. 1987. Supercritical extraction of toxic organics from soils. Industrial & Engineering Chemistry Research, 26, 261–268. Breinholt, B., Lauridsen, S. T., Daneshvar, B. and Jakobsen, J. 2000. Dose–response effects of lycopene on selected drug-metabolizing and antioxidant enzymes in the rat, Cancer Letters, 154, 201–210. Cadoni, E., Giorgi, M. R., Medda, E. and Poma, G. 2000. Supercritical CO2 extraction of lycopene and b-carotene from ripe tomatoes. Dyes and Pigments, 44, 27–32. Calvey, E. M. and Page, S. W. 1990. Apparent solubility threshold densities of substituted coumarins, Journal of Supercritical Fluids, 3, 115–120. Casas, L., Mantell, C., Rodríguez, M., Torres, A., Macías, F.A. and Martínez de la Ossa, E. 2007. Effect of the addition of cosolvent on the supercritical fluid extraction of bioactive compounds from Helianthus annuus L. Journal of Supercritical Fluids, 41, 43–49. Chandra, A. and Nair, M. G. 1997. Supercritical fluid carbon dioxide extraction of a- and b-carotene from carrot (Daucus carota L.). Phytochemical Analysis, 8, 244–246. Chang, C. J. and Randolph, A. D. 1989. Precipitation of microsize organic particles from supercritical fluids. AIChE Journal, 35(11), 1879–1882. Chao, R. R., Mulvaney, S. J. and Hanah, H. 1993. Effects of extraction and fractionation pressures on supercritical extraction of cholesterol from beef tallow, Journal of the American Oil Chemists’ Society, 70, 139–143. Ciurlia, L., Bleve, M. and Rescio, L. 2009. Supercritical carbon dioxide co-extraction of tomatoes (Lycopersicum esculentum L.) and hazelnuts (Corylus avellana L.): a new procedure in obtaining a source of natural lycopene. Journal of Supercritical Fluids, 49(3), 338–344 Cocero, M. J., González, S., Pérez, S. and Alonso, E. 2003. Supercritical extraction of unsaturated products. Degradation of b-carotene in supercritical extraction process, Journal of Supercritical Fluids, 19, 39–44. Clinton, S. K. 1998. Lycopene: chemistry, biology, and implications for human health and disease. Nutrition Reviews, 56(2), 35–51. Cygnarowicz, M. L., Maxwell, R. J. and Seider, W. D. 1990. Equilibrium solubilites of b-carotene in supercritical carbon dioxide. Fluid Phase Equilibria, 59, 57–71. Dewanto, V., Wu, X., Adom, K. K. and Liu, R. H. 2002. Thermal processing enhances the nutritional value of tomatoes by increasing total antioxidant activity. Journal of Agriculture and Food Chemistry, 50, 3010–3014. Ferreira, S. R. S. and Meireles, M. A. A. 2002. Modeling the supercritical fluid extraction of black pepper (Piper nigrum L.) essential oil. Journal of Food Engineering, 54, 263–269. Franca, L. F. and Meireles, M. A. A. 2000. Modeling the extraction of carotene and lipids from pressed palm oil (Elaes guineensis) fiber using supercritical CO2. Journal of Supercritical Fluids, 18, 35–47. Friedrich, J. P. and Pryde, E. H. 1984. Supercritical CO2 extraction of lipid-bearing materials and characterization of products, Journal of the American Oil Chemists Society, 61, 223. Gordon, H. T. and Bauernfeind, J. C. 1982. Carotenoids as food colorants. Critical Reviews in Food Science and Nutrition, 18, 59–97. Gouveia, L., Nobre, B. P., Marcelo, F. M., Mrejen, S., Cardoso, M. T., Palavra, A. F. and Mendes, R. L. 2006. Functional food oil coloured by pigments extracted from microalgae with supercritical CO2. Food Chemistry, 43, 2876–2878. Güçlü-Üstündag, Ö. and Temeli, F. 2005. Solubility behavior of ternary systems of lipids, cosolvents and supercritical carbon dioxide and processing aspects. Journal of Supercritical Fluids, 36, 1–15.
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642 Separation, extraction and concentration processes Hackett, M. M., Lee, J. H., Francis, D. and Schwartz, S. J. 2004. Thermal stability and isomerisation of lycopene in tomato oleoresins from different varieties. Journal of Food Science, 69, 536–541. Hadley, C. W., Clinton, S. K. and Schwartz, S. J. 2003. The consumption of processed tomato products enhances plasma lycopene concentrations in association with reduced lipoprotein sensitivity to oxidative damage. Journal of Nutrition, 133, 727–732. Hansen, B. N., Harvey, A. H., Coelho, J. A. P., Palavra, A. M. F. and Bruno, T. J. 2001. Solubility of capsaicin and b-carotene in supercritical carbon dioxide and in halocarbons. Journal of Chemical and Engineering Data, 46, 1054–1058. Hawthorne, S. B., Galy, A. B., Schmitt, V. O. and Miller, D. J. 1995. Effect of SFE flow rate on extraction rates: classifying sample extraction behavior. Analytical Chemistry, 67, 2723–2732. Hopper, M. L. and King, J. W. 1991. Enhanced supercritical fluid carbon dioxide extraction of pesticides from foods using palletized diatomaceous earth, Journal of the Association of Official Analytical Chemists, 7, 661–666. Johannsen, M. and Brunner, G. 1997. Solubilities of the fat-soluble vitamins A, D, E, and K in supercritical carbon dioxide. Journal of Chemical and Engineering Data, 42, 106–111. Joslin, C. G., Gray, C. G. and Goldman, S. 1996. Solubility in supercritical fluids from the virial equation of state, Molecular Physics, 89, 489–503. Kassama, L. S., Shi, J. and Mittal, G. S. 2008. Optimization of supercritical fluid extraction of lycopene from tomato skin with central composite rotatable design model, Separation and Purification Technology, 60, 278–284. Krichnavaruk, S., Shotipruk, A., Goto, M. and Pavasant, P. 2008. Supercritical carbon dioxide extraction of astaxanthin from Haematococcus pluvialis with vegetable oils as co-solvent. Bioresource Technology, 99, 5556–5560. Krinsky, N. I. and Rock, C. L. 1998. Carotenoids: chemistry, sources and physiology. In M. Sadler, S. Strain, B. Caballero (Eds.), Encyclopedia of Human Nutrition (pp. 304–314). Academic Press, London, UK. Lehotay, S. J. 1997. Supercritical fluid extraction of pesticides in foods. Journal of Chromatography A, 785, 289–312. Liu, L. H., Zabaras D., Bennett, L. E., Aguas, P. and Woonton, P. W. 2009. Effects of UV-C, red light and sun light on the carotenoid content and physical qualities of tomatoes during post-harvest storage. Food Chemistry, 115, 495–500. Luque de Castro, M. D., Valcarcel, M. and Tena, M. T. 1994. Analytical supercritical fluid extraction. Springer-Verlag, Berlin, Germany. Marentis, R. T. 1988. Steps to developing a commercial supercritical carbon dioxide processing plant. In Supercritical fluid extraction and chromatography. Charpentier, B. A. and Sevenants, M. R., Ed.; American Chemical Society, Symposium Series, 127–143. Marsili, R. and Callahan, D. 1993. Comparison of a liquid solvent extraction technique and supercritical fluid extraction for the determination of a- and b-carotene in vegetable. Journal of Chromatographic Science, 31, 422–428. Mascio, P. D., Kaiser, S. and Sies, H. 1989. Lycopene as the most efficient biological carotenoid singlet oxygen quencher. Archives of Biochemistry and Biophysics, 274, 532–538. Mayeaux, M., Xu, Z., King, J. M. and Prinyawiwatkul, W. 2006. Effects of cooking conditions on the lycopene content in tomatoes. Journal of Food Science, 71, 461–464. Mayer-Miebach, E., Behsnilian, D., Regier, M. and Schuchmann, H. P. 2005. Thermal processing of carrots: lycopene stability and isomerisation with regard to antioxidant potential. Food Research International, 38, 1103–1108. Mayer-Miebach, E. and Spiess, W. E. L. 2003. Influence of cold storage and blanching on the carotenoid content of Kintoki carrots. Journal of Food Engineering, 65, 211–213. © Woodhead Publishing Limited, 2010
Supercritical-fluid extraction of lycopene from tomatoes 643 Nagy, B. and Simándi, B. 2008. Effects of particle size distribution, moisture content, and initial oil content on the supercritical fluid extraction of paprika. Journal of Supercritical Fluids, 46(3), 293–298. Negre-Salvayre, A., Dousset, N., Ferretti, G., Bacchetti, T., Curatola, G. and Salvayre, R. 2006. Antioxidant and cytoprotective properties of high-density lipoproteins in vascular cells. Free Radical Biology and Medicine, 41, 1031–1040. O’Day, D. M. and Rosenau, J. R. 1982. Solvent extraction of carotenoids from alfalfa. Transactions of the ASAE (American Society of Agricultural Engineers), 25, 515– 519. Ollanketo, M., Hartonen, K., Riekkola, M. L., Holm, Y. and Hiltunen, R. 2001. Supercritical carbon dioxide extraction of lycopene in tomato skins. European Food Research and Technology, 212, 561–565. Olson, J. 1986. Carotenoid, vitamin A and cancer. Journal of Nutrition, 116, 1127– 1130. Palozza, P. 1998. Prooxidant actions of carotenoids in biologic systems. Nutrition Reviews, 56(9), 257–265. Papas, A. 1999. Diet and antioxidants status: In Antioxidant status, diet, nutrition and health. Ed. Andreas M. Papas. pp. 89–106. CRC Press, New York. Pol, J., Hyotylainen, T., Ranta-Aho, O. and Riekkola, M-L. 2004. Determination of lycopene in food by on-line SFE coupled to HPLC using a single monolithic column for trapping and separation. Journal of Chromatography A, 1052, 25–31. Pourmortazavi, S. M. and Hajimirsadeghi, S. S. 2007. Supercritical fluid extraction in plant essential and volatile oil analysis. Journal of Chromatograph A, 1163, 2–24. Rao, A. V. and Agarwal, S. 1998. Bioavailability and in vivo antioxidant properties of lycopene from tomato products and their possible role in the prevention of cancer. Nutrition and Cancer. 31, 199–203. Rao, A. V. and Agarwal, S. 1999. Role of lycopene as antioxidant carotenoid in the prevention of chronic diseases: a review. Nutrition Research, 19(2), 305–32. Rao, A.V. and Rao, L.G. 2007. Carotenoids and human health. Pharmacological Research, 55(3), 207–216. Rozzi, N. L., Singh, R. K., Vierling, R. A. and Watkins, B. A. 2002. Supercritical fluid extraction of lycopene from tomato processing by-products. Journal of Agriculture and Food Chemistry, 50, 2638–2643. Sabio, E., Lozano, M., Montero de Espinosa, V., Mendes, R. L., Pereira, A. P., Palavra, A. F. and Coelho, J. A. 2003. Lycopene and b-carotene extraction from tomato processing waste using supercritical CO2. Industrial & Engineering Chemistry Research, 42, 6641–6646. Sakaki, K. 1992. Solubility of b-carotene in dense carbon dioxide and nitrous oxide from 308 to 323 K and from 9.6 to 30 MPa. Journal of Chemical and Engineering Data, 37, 249–251. Seybold, C., Frohlich, K., Bitsch, R., Otto, K. and Bohm, V. 2004. Changes in contents of carotenoids and vitamin E during tomato processing. Journal of Agriculture and Food Chemistry, 52, 7005–7010. Shi, J. 2002. Lycopene: biochemistry and functionality. Food Science and Biotechnology, 11(5), 574–581. Shi, J., Bryan, M., Le Maguer, M. and Kakuda, Y. 2002a. Kinetics of lycopene degradation in tomato puree with heat and light irradiation treatment. Journal of Food Processing Engineering, 25, 485–498. Shi, J., Dai, Y., Kakuda, Y., Mittal, G. and Xue, J. 2008. Effect of heating and light irradiation on the stability of lycopene in tomato purée. Food Control, 19(5), 514–520. Shi, J., Kakuda, Y., Zhou, X., Mittal, G. and Pan, Q. 2007a. Correlation of mass transfer coefficient in the extraction of plant oil in a fixed bed for supercritical CO2. Journal of Food Engineering, 78, 33–40. Shi, J., Kakuda Y. and Yeung, D. 2004a. Antioxidative properties of lycopene and other carotenoids: synergistic effects. BioFactors, 21, 203–210. © Woodhead Publishing Limited, 2010
644 Separation, extraction and concentration processes Shi, J., Kassama, L. and Kakuda, Y. 2006. Supercritical fluid technology for extraction of bioactive components. Functional food ingredients and nutraceuticals: processing technology, Ed. J. Shi. CRC Press, USA. p. 45–73. Shi, J., Khatri, M., Xue, J., Mittal, G., Ma, Y. and Li, D. 2009a. Solubility of lycopene in supercritical CO2 fluid affected by temperature and pressure. Separation and Purification Technology, 66, 322–328. Shi, J., Le Maguer, M. and Kakuda, Y. 1999a. Lycopene degradation and isomerization in tomato dehydration. Food Research International, 32, 15–21. Shi, J., Le Maguer, M. and Wang, S. 1999b. Chemical composition of tomatoes affected by maturity and fertility practices. Journal of Food Quality, 22, 147–156. Shi, J., MacNaughton, L., Kakuda, Y., Bettger, W., Yeung, D. and Jiang, Y. 2004c. Bioavailability of lycopene from tomato products. Journal of Food Science and Nutrition, 9, 98–106. Shi, J., Mittal, G., Kim, E. and Xue, S. 2007b. Solubility of carotenoids in supercritical CO2. Food Review International, 23, 341–371. Shi, J., Qu, Q., Kakuda, Y., Xue, J., Jiang, Y., Koide, S. and Shim, Y. 2007c. Investigation of the antioxidant and synergistic effect of lycopene and vitamin E on the AMVN induced oxidation of LAME. Journal of Food Composition and Analysis, 20(6), 603–608. Shi, J., Qu, Q., Kakuda, Y. and Yeung, D. 2004c. Stability and bioavailability and synergistic effect of antioxidant capacity of lycopene with other antioxidants. Critical Reviews in Food Science and Nutrition, 44, 559–573. Shi, J., Wu, Y., Bryan, M. and Le Maguer, M. 2002b. Oxidation and isomerization of lycopene under thermal treatment and light irradiation in food processing. Nutraceuticals and Foods, 7(2), 179–183. Shi, J., Yi, C., Xue, J., Jiang, Y., Ma, Y. and Li, D. 2009b. Effects of modifier on lycopene extract profile from tomato skin using supercritical-CO2 fluid. Journal of Food Engineering, 93, 431–436. Shi, J., Zhou, X. and Kassama, L. 2007d. Correlation of mass transfer coefficient in separation process with supercritical CO2. Drying Technology International, 25, 335–339. Shi, J. and Le Maguer, M. 2000. Lycopene in tomatoes: chemical and physical properties affected by food processing. Critical Reviews in Food Science and Nutrition CRC, 40(1), 1–41. Shi, J. and Zhou, X. 2006. Solubility property of bioactive components on recovery yield in separation process by supercritical fluid. Functional food ingredients and nutraceuticals: processing technology, Ed. J. Shi. CRC Press, USA. p. 3–43. Shixian, Q., Dai, Y., Kakuda, Y., Shi, J., Mittal, G., Yeung, D. and Jiang,Y. 2005. Synergistic antioxidative effects of lycopene with other bioactive compounds. Food Review International, 21, 295–311. Simandi, B., Kristo, S. T., Kery, A., Selmeczi, L. K., Kmecz, I. and Kemeny, S. 2002. Supercritical fluid extraction of dandelion leaves, Journal of Supercritical Fluids, 23, 135–142. Sovová, H., Stateva, R. P. and Galushko, A. A. 2001. Solubility of b-carotene in supercritical CO2 and the effect of entrainers. Journal of Supercritical Fluids, 21, 195–203. Spanos, G. A., Chen, H. and Schwartz, S. J. 1993. Supercritical CO2 extraction of b-carotene from sweet potatoes. Journal of Food Science, 58(4), 817–820. Stahl, W. and Sies, H. 1992. Uptake of lycopene and its geometrical isomers is greater from heat-processed than from unprocessed tomato juice in humans. Journal of Nutrition, 122, 2161–2166. Stahl, W. and Sies, H. 1996. Perspectives in biochemistry and biophysics. Archives of Biochemistry and Biophysics, 336 (1), 1–9. Subra, P., Castellani, S., Ksibi, H. and Garrabos, Y. 1997. Contribution to the determination of the solubility of b-carotene in supercritical carbon dioxide and nitrous oxide, experimental data and modelling. Fluid Phase Equilibria, 131, 269–286. © Woodhead Publishing Limited, 2010
Supercritical-fluid extraction of lycopene from tomatoes 645 Suganuma, H. and Inakuma, T. 1999. Protective effect of dietary tomato against endothelial dysfunction in hypercholesterolemic mice. Bioscience, Biotechnology, and Biochemistry, 63, 78–82. Sun, M. and Temelli, F. 2006. Supercritical carbon dioxide extractions of carotenoids from carrot using canola oil as a continuous co-solvent. Journal of Supercritical Fluids, 37, 397–408. Tonthubthimthong, P., Chuapraset, S., Douglas, P. and Luewisutthicat, W. 2001. Supercritical CO2 extraction of nimbin from neem seeds – an experimental study, Journal of Food Engineering, 47, 289–293. Toor, R. K. and Savage, G. P. 2006. Effect of semi-drying on the antioxidant components of tomatoes. Food Chemistry, 94, 90–97. Topal, U., Sasaki, M., Goto, M. and Hayakawa, K. 2006. Extraction of lycopene from tomato skin with supercritical carbon dioxide: Effect of operating conditions and solubility analysis. Journal of Agriculture and Food Chemistry, 54, 5604–5610. Tsai, C., Lin, H. and Lee, M. 2007. Solubility of disperse yellow 54 in supercritical carbon dioxide with or without cosolvent, Fluid Phase Equilibria, 260, 287–294. Vági, E., Simándi, B., Vásárhelyiné, K. P., Daood, H., Kéry, Á., Doleschall, F. and Nagy, B. 2007. Supercritical carbon dioxide extraction of carotenoids, tocopherols and sitosterols from industrial tomato by-products. Journal of Supercritical Fluid, 40, 218–226. Vasapollo, G., Longo, L., Rescio, L. and Ciurlia, L. 2004. Innovative supercritical CO 2 extraction of lycopene from tomato in the presence of vegetable oil as co-solvent. Journal of Supercritical Fluids, 29, 87–96. Vaughn, K. L. S., Clausen, E. C., King, J. W., Howard, L. R. and Carrier, D. J. 2008. Extraction conditions affecting supercritical fluid extraction (SFE) of lycopene from watermelon. Bioresource Technology, 99(16), 7835–7841. Vega, P. J., Balaban, M. O., Sims, C. A., Keefe, S. F. and Cornell, J. A. 1966. Supercritical carbon dioxide extraction efficiency for carotenes from carrots by RSM, Journal of Food Science, 61, 757–765. Walsh, J. M., Ikonomou, G. D. and Donohue, M. D. 1987. Supercritical phase behavior: the entrainer effect. Fluid Phase Equilibria, 33, 295–314. Willcox, J. K., Catignani, G. and Lazarus, S. 2003. Tomatoes and cardiovascular health. Critical Reviews in Food Science and Nutrition, 43, 1–18. Wu, D. and Meydini, S. N. 1999. Antioxidants and immune function: In Antioxidant status, diet, nutrition and health. ed. Andreas M. Papas. pp. 371–400. CRC Press, New York. Xianquan, S., Shi, J., Kakuda, Y. and Jiang, Y. 2005. Stability of lycopene during food processing and storage. Journal of Medicinal Food, 8(4), 413–422. Yi, C., Shi, J., Xue, J., Jiang, Y. and Li, D. 2009. Effects of supercritical fluid extraction parameters on lycopene yield and antioxidant activity. Food Chemistry, 113(4), 1088–1094. Ziémons, E., Goffin, E., Lejeune, R., Proença da Cunha, A., Angenot, L. and Thunus, L. 2005. Supercritical carbon dioxide extraction of tagitinin C from Tithonia diversifolia. Journal of Supercritical Fluids, 33, 53–59.
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646 Separation, extraction and concentration processes
© Woodhead Publishing Limited, 2010
Index 647
Index
a-lactalbumin, 125 properties and applications, 453–4 purification, 461–2 chromatographic techniques, 461 membrane filtration, 461 other techniques, 461–2 accelerated solvent extraction (ASE), 40 see also pressurised liquid extraction acetonitrile, 48 Achiever press, 411, 412 active membranes, 318, 320 activity coefficient models, 17 adalimumab, 168 adsorbents nutraceutical separation, 148–72 membrane adsorbers, 169–72 mesoporous molecular sieves, 163–6 molecular imprinted polymers, 149–53 organic monoliths, 153–9 peptide affinity ligands and phage display methods, 166–9 stimuli-responsive resins, 159–63 adsorptive membrane separation, 458 aerobic membrane bioreactor, 191 affinity chromatography, 113–14, 131, 166–7, 468–9, 472–3 affinity separations, 163
air gap membrane distillation (AGMD), 189, 266 Amberlite, 124 Amberlite XAD-16, 54 anaerobic membrane bioreactor, 191 analytical separation techniques, 238 anthocyanins, 51, 94, 100 antibiotics, 57 antioxidants activity in food systems, 507–11 in vitro methods for pure compounds and extracts, 509–10 lipid oxidation mechanism, 508 biological activities, 521–6 plant food extracts, 522–3 definition, 507 extraction from plant foods, 506–67 future trends, 567 extraction from plant foods and residues, 526–56 conventional solvent extraction, 526–31 conventional solvent extraction examples, 528–30 enzyme-aided extraction, 531–8 effect of enzyme concentration on yield of extractable compounds, 535
© Woodhead Publishing Limited, 2010
648 Index
high pressure systems, 541–6 other extraction technologies, 553–5 other novel extraction technologies, 552–6 pressurised liquid extraction, 538–7 supercritical-fluid extraction, 547–52 extraction processes and purification integration, 556–67 integration of extraction processes and purification membrane distillation and osmotic evaporation, 560 membrane fractionation and resin purification, 566 protease hydrolysis and membrane fractionation of peptides, 562 protein hydrolysate production, 563 supercritical carbon dioxide and ultra- and nanofiltration, 559 natural compounds and major sources, 511–21 fruits, vegetables and cereals phenolic composition, 516–18 natural compounds, 511–16 plants phenolic contents, 513–14 residual sources proposed for manufacture, 518 solid agro-industrial by-products chemical composition, 519–20 sources, 516–21 aroma, 230 aroma compounds recovery by pervaporation and applications in food industry, 230–9 analytical separation technique, 238 dairy aromas recovery, 238 feed temperature effect on methyl butyrate solution selectivity, 234 feed temperature effect on total mass flux for methyl butyrate solution, 233 fruit juice aromas recovery, 231–6
other applications, 239 tea, cocoa and coffee aromas recovery, 237–8 wine aromas recovery, 236–7 Arrhenius equation, 235 ASE 150, 61 ASE 350, 53, 61 ASPEN Chrom, 118 astaxanthin, 299–300, 301, 302, 306–7 auxiliary finings, 436–7 avidin, 373–4 b-carotene, 638 b-galactosidase treatment, 192 b-lactoglobulin, 125, 131 properties and applications, 451–3 purification, 454–60 chromatographic techniques, 458–60 membrane chromatography, 460 membrane filtration, 460 selective precipitation, 454–5, 458 b-ovomucin, 599 backflushing, 195–6, 443 bactofugation, 349 Bactofuge, 349 beer, 122–3 filtration, 443–6 betalaines, 79 betulin, 49 bioactive peptides, 363–6, 451 bioproducts separation by colloidal gas aphrons flotation, 284–309 CGA applications, 293–307 CGA properties and applications, 285–93 future trends, 308–9 industrial application feasibility, 307–8 bipolar membrane, 212 bovine serum albumin properties and applications, 454 purification, 462 BPM see bipolar membrane brewery products characteristics, 431–2 technology and raw materials selection, 432 brewing
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Index 649
bulk beer filtration by membranes, 443–6 membrane fouling mechanisms, 443 filter aid filtration and applications, 437–41 configuration, 438 kieselguhr or diatomaceous earth, 437 perlite, 440 particle size and concentration distribution beer samples before and after filtration, 440 pre-filter beer, 439 separation technologies, 430–48 beer fining agents, 436–7 brewery products characteristics, 431–2 cleaning agents recovery, 446 dissolved gas control by membrane technology, 446–7 future trends, 447–8 regeneratable and reusable filter aids and applications, 441–2 technology and raw materials selection, 432 wort production in brewhouse, 433–4 yeast flocculation and applications, 435–6 whirlpools and applications, 434–5 whirlpool for trub separation, 434 bulk polymerisation methods, 152 Candida antarctica lipase B (CALB), 330 canola oil, 633, 638 carbohydrates, 515–16, 537 carnosic acid, 540 carotenoids, 526 carousel extractor, 420 cartridge filters, 441 casein micelles, 344, 450 microfiltration, 355–6 casein-whey protein co-precipitates, 351 caseins, 344 fractionation, 361 cation-exchange chromatography (CEC), 126, 127, 128, 464 cation exchange resins, 120
cavitation, 93 cellulose nitrate, 220 centrifugal evaporators, 245 centrifugal partition chromatography (CPC), 487, 488 centrifugation, 348–9 Ceratocystis moniliformis, 239 cereals, 520 chiral separation, 132 chromatography, 149 see also specific chromatographic methods operating modes, 116–17 Clapeyron’s law, 260 cleaning-in-place technology, 446 closed-wall extruders, 400–1 CM Sephadex, 124 co-solvent, 631–4 cold energy, 387 colloidal gas aphrons (CGA), 469 applications in recovery of valueadded food products separation process, 294–5 applications in value-added food products recovery, 293–307 bioproducts separation by flotation, 284–309 feasibility of industrial applications, 307–8 future trends, 308–9 characteristics, 289–90 generated by ionic and non-ionic surfactants, 291 generation, 291–3 formation, 291–2 impeller with laboratory mixer, 292 surfactants, 292–3 properties and applications in food industry, 285–93 recovery of particles and mechanism separation, 299–307 applications in flotation columnsetup, 303–5 astaxanthin recovery and mechanism of separation, 299–300, 301 recovery of particles, 299 scale-up of astaxanthin separation, 302
© Woodhead Publishing Limited, 2010
650 Index
scale-up with a flotation column, 300, 302–7 surfactants used and outcomes on astaxanthin recovery, 301 recovery of soluble compounds and mechanism of separation, 296–9 bioactive components from plant extracts, 298–9 proteins, 296–8 small-scale procedure for protein separation, 297 structure, 285–9 proposed structure, 286 studies for characterisation, 287–8 colloidal instability, 431 combine stabilisation system (CSS), 123 composite membranes, 228 Comsol, 118 conalbumin, 598 concentration methods omega-3 fatty acids, 483–502 chromatographic methods, 486–8 distillation methods, 492–5 enzymatic methods, 495–8 integrated methods, 498–501 low temperature fractional crystallisation, 488–90 supercritical-fluid extraction, 490–2 urea adduction, 484–6 concentration polarisation, 226, 264, 269, 271 conductivity disintegration index Z, 74, 83–4 continuous membrane chromatography reactor system (CMCRS), 136 continuous recycle membrane bioreactors (CRMBR), 316, 324, 328 continuous separation technology, 149 continuous stirred tank reactor (CSTR), 323 convective mass transfer, 157 conventional heating, 97–8 counterions, 203 covalent imprinting method, 150 cranberry fruits, 211 cross flow microfiltration, 443, 445 crosslinked enzyme aggregates (CLEAs), 320 cryoconcentration, 245–6
CSEP see continuous separation technology cyclodextrin glucosyl transferase (CGTase), 328 dairy industry dairy products compositions, 343–7 milk, 343–6 whey, 346–7 fractionation of proteins and peptides, 360–6 caseins, 361 lysozyme extraction, 364 marketed bioactive peptides, 365 milk protein functionality, 360 peptides, 363–6 serum proteins, 361–3 milk proteins standardisation and concentration, 351–4 cheese production and MMV process, 352 milk separation techniques, 347–51 bacterial removal, 348–51 microfiltration, 350 skimming and fractionation of fat globules, 347–8 uniform transmembrane pressure system, 350 separation techniques for wheys and derivatives in cheese production, 357–60 concentration and demineralisation, 358–60 demineralisation of whey, 359 serum proteins concentration, 357–8 separation technologies, 341–69 effluents and technical fluids, 366–8 future trends, 368–9 whole casein isolation, 354–6 isoelectric precipitation and rennet coagulation, 354–5 microfiltration, 355–6 Primin process, 356 dairy nutraceuticals, 450–74 acidic whey protein components, 451–4 a-lactalbumin properties and applications, 453–4
© Woodhead Publishing Limited, 2010
Index 651
b-lactoglobulin properties and applications, 451–3 bovine serum albumin properties and applications, 454 milk proteins and their bioactivities, 452 basic proteins, 462–3 lactoferrin, 462–3 lactoperoxidase, 463 lysozyme, 463 immunoglobulins, 470–1 immunoglobulin A, 470 immunoglobulin G, 470 immunoglobulin M, 471 purification methods, 450–74 acidic whey proteins, 454–62 basic whey proteins, 463–70 future trends, 473–4 immunoglobulins, 471–3 dealcoholisation, 236 degree of damage, 74 desalination, 272 Desmet Ballestra press, 411 Desmet LM extractor, 418, 420 diffusion, 238 diffusion coefficient, 224 Dionium, 61 direct contact membrane distillation (DCMD), 187, 266, 269, 273 direct osmosis membrane-based concentration advances, 248–50 applications, 249–50 operating conditions effect, 248–9 process fundamentals, 248 displacement chromatography, 117 distillation, 492–5 Donnan exclusion, 203 Dox-Hivex extruders, 402–3 dry extrusion, 402 Duolite C-464, 124, 609 Dupps screw press, 411 Duvis press, 413–15, 414 dynamic PFE, 42 edible oil, 633 effluents, 366–8 egg products bioactive peptides derived from egg proteins, 603
structural properties, 606 compositions, 369–71 egg white and whole egg concentration and stabilisation, 370–1 egg white proteins, 370 hen egg characteristics, 369–70 egg white proteins composition, 596 lysozyme, 599 minor egg white proteins, 599 ovalbumin, 597–8 ovomucin, 598–9 ovomucoid, 598 ovotransferrin, 598 physicochemical characteristics, 597 egg white proteins extraction, 371–4 avidin, 373–4 lysozyme, 372–3 ovotransferrin, 373 egg yolk proteins, 599–601 composition, 596 lipovitellenin, 601 lipovitellins, 601 livetins, 601 phosvitin, 600 hen eggs chemical composition and major components, 596 proteins and peptides composition and physicochemical characteristics, 597–601 fractionation for nutraceutical applications, 595–613 proteins and peptides biological activities, 601–5 ACE inhibitor, antihypertensive and vasorelaxing activities, 602 antimicrobial activity, 604–5 antioxidant activity, 605 separation technologies, 369–76 future trends, 375–6 yolk components extraction, 374–5 technologies for egg proteins and peptides fractionation, 605–12 chromatographic methods, 608–10 membrane processes, 610–12 other separation methods, 612 precipitation, 607–8
© Woodhead Publishing Limited, 2010
652 Index
technological approaches and processes, 606 elastomers, 228 electrical theory, 91 electro-osmosis, 85 electrochemical coagulation, 204 electrodialysis, 202, 265 electrodialysis with bipolar membranes (EDBPM), 212 electrodialysis with filtration membranes (EDFM), 206, 211 electrodialysis with ion-exchange membranes, 206 electrodialytic phenomena applications, 204–13 associated membrane technologies and applications, 202–14 electrodialysis with bipolar membranes, 212–13 passion fruit juices deacidification, 212–13 11S–7S fractionation, 213 electrodialysis with filtration membranes, 206, 211–12 anticancer peptidic fraction production, 211 cranberry juice antioxidant enrichment, 211–12 electrodialysis with ion-exchange membranes, 206 main applications, 207–10 electrolysis with membranes, 204–6 electrodialytic phenomena and membrane configurations, 205 lipid stability enhancement of omega-3 enriched milk, 204, 206 milk protein electrochemical coagulation, 204 future trends, 213–14 principles, 203 electrode phenomena, 203 membrane phenomena, 203 electrofiltration, 197 electromembrane filtration, 112 electropermeabilisation, 73, 83 electroplasmolysis, 73 electroporation, 73 electroreduction, 204 elution chromatography, 117
enhanced solvent extraction (ESE), 40 enrichment factors, 225, 235, 238 enzymatic hydrolysis, 499–500 enzyme-aided extraction, 531–8 plant food antioxidants, 532–4 ethanol, 632, 638 European contract Enhance QLK5 CT 199901442, 127 expanded bed adsorption (EBA), 111, 135, 467, 471–2 explosives, 57 extraction, 625 extrusion, 399–403 extrusion-screw press system, 410 falling film evaporators, 245 faradaic reactions, 203 fast-PSE, 61 ferulic acid, 538 Fick’s law, 223, 254 filter aid, 441–2 filtration, 437–41 finings, 436–7 fixed-bed chromatography, 122 flash desolventiser, 423 flavanols, 123 flavonoids, 512, 524, 525, 526 flavour instability, 431 flavour stability, 431 flotation, 285, 300 foam fractionation, 285, 296, 298 focused mode see single-mode food industry aroma compounds recovery by pervaporation, 219–39 analytical separation technique, 238 dairy aromas, 238 fruit juice aromas, 231–6 membranes selection, 227–9 principles, 220–1 tea, cocoa and coffee aromas, 237–8 transport mechanism, 221–7 wine aromas, 236–7 electrodialytic phenomena and membrane technologies, 202–14 applications, 204–13 future trends, 213–14 principles, 203
© Woodhead Publishing Limited, 2010
Index 653
membrane technologies application, 180–97 dairy industry, 182–5 fruit juice industry, 185–90 future trends, 195–7 new applications for saccharides concentration and fractionation, 191–5 theoretical fundamentals, 181–2 wastewater treatment, 190–1 physically assisted extraction methods, 71–102 high-voltage electrical discharges, 86–90 microwave-assisted extraction, 96–100 ohmic heating assisted extractions, 83–6 physical treatments combination, 100–2 pulsed electric field assisted extractions, 72–83 ultrasound-assisted extraction, 90–6 process chromatography, 109–37 applications, 118–28 basic principles, 113–18 future trends, 135–7 process control, 135 recent developments, 128–35 supercritical fluid extraction principles and applications, 3–36 cycle processes for extraction, 21–6 fundamentals of thermodynamics, 8–20 liquids extraction, 30–2 solids extraction, 26–30 food ingredients, 314–31 fouling, 182–3, 191 fractional crystallisation, 488–90 fractional distillation, 492–3 fractionation, 467–8 available technologies for egg proteins and peptides, 605–12 chromatographic methods, 608–10 membrane processes, 610–12 other separation methods, 612 precipitation, 607–8
technological approaches and processes, 606 egg proteins/peptides for nutraceutical applications, 595–613 biological activities, 601–5 composition and physicochemical characteristics, 597–601 proteins and peptides in dairy industry, 360–6 11S–7S fractionation, 213 free cells membrane bioreactors (FCMBR), 315, 316–18 free enzymes membrane bioreactors (FEMBR), 315, 316–18, 323, 327–9 freeze concentration, 245 frontal chromatography, 117 fructose, 120 fruit juices aroma recovery by pervaporation, 231–6 concentrate production, 386–94 concentration by evaporation, 386–7 concentration by freezing, 387 concentration by membrane separation processes, 387–94 Furmint must concentration on pilot-scale reverse osmosis plant, 390 grape juice concentration by complex membrane processes, 388 Hungarian musts Kèkfrancos and Furmint, 391 must concentration total cost changes, 394 permeate flux changing during clarification, 389 designing separation processes to optimise product quality, 383–6 juice clarification, 385–6 juice extraction by pressing, 383–5 juice extraction wing water as solvent, 385 natural fruit juice and fruit juice production, 384 separation technologies, 381–94 foods/fluids characteristics in product sector, 382–3
© Woodhead Publishing Limited, 2010
654 Index fugacity, 12, 18 fugacity coefficients, 13 gallic acid, 299 gas hold-up, 289 gas membrane extraction see osmotic distillation gas-phase fugacity coefficient, 13 gel filtration, 134 gel-permeation chromatography, 608–9 glassy membrane, 228 glassy polymers, 227 glycomacropeptide, 366 GP Membralox membranes, 349 gPROMS, 118 green technology, 621 half concentrate, 388–90 Harburg-Freunberger press, 411–13, 412 hazelnut oil, 633 heat transfer equations, 255 high-diffusion liquids, 40 high frequency backflushing, 195–6 high gravity brewing, 447 high hydrostatic pressure (HHP), 552, 556 high liquid entry pressure, 269 high-performance tangential flow filtration (HPTFF), 461 high-pressure liquid extraction see pressurised liquid extraction high-speed countercurrent chromatography, 487–8 high-voltage electrical discharges (HVED), 86–90 experimental setup and voltage and current graphs, 87 solute extraction, 87–90 grape pomace, 89–90 oil, 88–9 treatment temperature effect on polyphenols final yield, 89 yeast cells, 90 underwater HVED principles effect, 86–7 electric discharge type in aqueous solution, 88 Hildebrand solubility parameter, 45 Hiplex process, 412 Hoechst-Celanese Liqui-Cel membrane contractor, 392
hollow-fibre membranes, 256 horizontal cookers, 398–9 hot ball model, 46 hot spot theory, 91 human monoclonal antibody in rheumatoid arthritis (HUMIRA), 168 hydrophilic pervaporation, 228 hydrophilic polymer films, 257 hydrophobic interaction chromatography, 113, 161–2 hydrophobic polymers, 256 immersed membrane bioreactor (iMBR), 317, 324 immobilised enzyme membrane reactor (IEMBR), 319–22, 323, 329–30 monophasic and biphasic reactors, 319 immunoglobulin A (IgA), 125, 470 immunoglobulin G (IgG), 125, 470 immunoglobulin M, 471 immunoglobulin Y, 607 immunoglobulins, 158, 363, 451, 470–1 purification methods, 471–3 affinity chromatography, 472–3 membrane chromatography, 472 protein-A mimetic affinity peptides, 473 in-process pasteurisation, 431 Insta-Pro extruder, 413 Insta-Pro screw press, 414 inulin, 119 ion exchange chromatography (IEC), 111, 113, 358, 362, 364, 372, 609 immunoglobulins purification, 471–2 lactoferrin and lactoperoxidase purification, 464, 467–8 stimuli-responsive resins applications, 160–1 Irganox 1076, 43–4 isenthalpic throttling, 21, 22–3 isinglass, 436–7 isocratic elution, 117 isoelectric precipitation, 354–5 Isoflux membrane, 349 isothermal membrane distillation see osmotic distillation juice clarification, 385–6
© Woodhead Publishing Limited, 2010
Index 655 juice extraction, 383 pressing, 383–5 wing water as solvent, 385 juice yield, 384 kieselguhr, 437, 439–40, 441–2, 445 kieselguhr filtration, 367 Knudsen diffusion, 254, 270 l-carrageenan, 612 lactalbumin, 357 lactoferrin, 126–7, 451, 462–3 purification, 464, 467–9 affinity chromatography, 468–9 ion-exchange chromatography, 464, 467–8 non-chromatographic methods, 469 other chromatographic methods, 469 lactoperoxidase, 126–7, 463 purification, 464, 467–9 affinity chromatography, 468–9 ion-exchange chromatography, 464, 467–8 non-chromatographic methods, 469 other chromatographic methods, 469 lactose, 451 Laplace equation, 270 lauter tun, 433 life cycle assessment (LCA), 39 lipases, 495–8 lipid peroxidation, 507, 508 lipid transfer protein (LTP), 128 lipovitellenin, 601 lipovitellins, 601 Lipozyme, 497 Liqui-Cel Extra-Flow, 257–8 Liqui-Cel membrane modules, 262 liquid-crystal templating, 164 livetins, 601 low-gradient plasmolysis, 84–5 lower critical solution temperature (LCST), 159 lycopene, 94, 99–100, 525 effect of co-solvent and modifiers on extraction, 631–4 binary and ternary modifiers, 633–4 edible oil, 633
ethanol, 632 modifiers on ratio of all-trans to cis-isomers, 634 water, 632–3 effects of pressure and temperature on antioxidant activity, 628–31 lycopene yield and antioxidant activity changes, 629 factors affecting extraction yield, 623–8 carbon dioxide flow rate, 625–6 operating time, 626 particle size, 627–8 pressure, 624 raw materials moisture content, 626–7 temperature, 624–5 factors affecting solubility in supercritical fluids, 635–8 co-solvents and modifiers, 637–8 pressure, 635–7 temperature, 637 solubility in supercritical fluids, 634–8 solubility isotherms, 636 structure, 620 supercritical-fluid extraction from tomatoes, 619–40 all-trans and cis-isomers contents and ratio in extract, 631 effects of modifiers on ratios of all-trans and cis-lycopene, 634 future trends, 639–40 HPLC chromatograms of the extract, 630 process, 622–3 lysozyme, 372–3, 463, 599, 604 purification, 469–70 Maillard polymers, 52 Maillard reaction, 52 Maillard reaction products (MRP), 515 marc desolventiser, 422 Marx Generator design, 81 mash filter technology, 433–4 mash tun, 433 mass separating agent, 25 mass transfer data, 26 mass transfer models, 5 mass transfer resistance, 253
© Woodhead Publishing Limited, 2010
656 Index Maubois, Mocquot and Vassal process, 351, 352, 353 MCM-41, 165, 166 MCM-48, 165, 166 membrane, 227 see also specific membrane membrane adsorbers, 169–72 advantages and disadvantages, 169–70 application, 170–1 affinity separations, 171 ion-exchange separations, 170–1 micro-organism removal, 171 commercialisation, 172 configurations, 170 functionalisation, 170 membrane-based concentration advances in food and beverage industries, 244–78 conventional evaporation and membrane concentration techniques key factors, 275 conventional technologies, 245–8 cryoconcentration, 245–6 thermal evaporation, 245 direct osmosis, 248–50 applications, 249–50 operating conditions effect, 248–9 process fundamentals, 248 membrane contactors, 250–75 membrane and osmotic distillation schematic representation, 251 osmotic and membrane distillation coupled operation, 273–5 MD and OD main advantages and disadvantages, 274 membrane distillation, 265–73 applications, 272–3 configuration types, 267 configurations, 266–7 mass transfer and polarisation phenomena, 268–9 MD membranes, 269–71 operating parameters effect on MD fluxes, 271–2 process fundamentals, 265–6 nomenclature, 276–8 osmotic distillation, 250–65 applications, 261–5 heat transfer, 255–6 mass transfer aspects, 252–65
membranes and modules, 256–8, 259 operating conditions effect on the OD flux, 258, 260–1 process fundamentals, 250–2 pressure-driven membrane processes, 246–8 commercial application, 247 membrane bioreactors (MBRs), 181 applications in food industries, 322–31 enzymatic MBRs, 325–6 food and beverage processing, 322–4 food ingredients production, 324–31 whole cell MBRs, 327 free enzymes or free cells membrane bioreactors, 316–18 continuous recycle membrane reactor, 316 submerged or immersed membrane bioreactor, 317 immobilised enzyme membrane reactor, 319–22 active membrane preparation and drawbacks, 320 monophasic and biphasic reactors, 319 production of food ingredients, 314–31 future trends, 331 membrane chromatography, 169–72, 460, 472 membrane distillation, 187–90, 265–73, 391, 558, 560 applications, 272–3 configurations, 266–7 types, 267 defined, 187 mass transfer and polarisation phenomena, 268–9 MD membranes, 269–71 operating parameters effect on MD fluxes, 271–2 process fundamentals, 265–6 membrane evaporation see osmotic distillation membrane fractionation, 558, 565–7 membrane ion-exchange chromatography, 171
© Woodhead Publishing Limited, 2010
Index 657 membrane molecular weight cut off (MWCO), 316 membrane osmotic distillation (MOD), 274 membrane separation process, 25–6, 368–9 membrane technologies application in food technology, 180–97 theoretical fundamentals, 181–2 dairy industry, 182–5 fouling, 182–3 fouling reduction advances, 184–5 membrane cleaning, 183–4 permeate flux vs transmembrane pressure, 183 fruit juice industry, 185–90 membrane distillation, 187–90 processes for concentration of juices, 188 reverse osmosis, 185–7 utilisation examples, 186 future trends, 195–7 backflushing influence on permeate flux vs time, 196 electrofiltration, 197 high frequency backflushing, 195–6 vibrating membrane module, 196 new applications for saccharides concentration and fractionation, 191–5 oligosaccharides ultrafiltration and nanofiltration membranes, 193 operational modes, 194–5 separation set-up design, 192 wastewater treatment, 190–1 aerobic membrane bioreactor, 191 anaerobic membrane bioreactor, 191 mesoporous molecular sieves, 163–6 advantages and disadvantages, 165 applications, 165–6 protein stabilisation, 166 size-based separations, 165 commercialisation, 166 liquid-crystal templating technique, 164 MCM-41 molecular sieve model, 164 synthesis, 164
methanol, 48, 638 microfiltration, 189, 348, 349, 350, 351, 355–6, 389, 610–12 microfiltration membranes, 246 microporosity, 270 microsieves, 445–6 microtitration, 130 microwave-assisted extraction (MAE), 96–100, 552 main applications, 98–100 food ingredients, 100 principles, 96–8 microwave vs conventional extraction, 97 reactors, 98 laboratory and industrial reactors, 99 safety considerations, 100 microwaves, 96 milk, 343–6 components physicochemical equilibria, 346 cows’ milk composition and association state, 343 cows’ milk mineral composition, 346 protein characteristics and biological functions, 345 proteins standardisation and concentration, 351–4 separation techniques, 347–51 bacterial removal, 348–51 microfiltration, 350 skimming and fractionation of fat globules, 347–8 uniform transmembrane pressure system, 350 mixed matrix membranes (MMMs), 131 MMV process see Maubois, Mocquot and Vassal process mobile crystalline materials (MCMs), 163 modifiers, 548 effect on lycopene extraction, 631–4 molasses, 122 molecular distillation, 493 molecular imprinted polymers (MIP), 149–53, 154 advantages and disadvantages, 152 applications, 152–3 commercialisation, 153 HPLC chromatograms, 154
© Woodhead Publishing Limited, 2010
658 Index preparation method, 151 synthesis, 150–2 mono-mode see single-mode monoliths, 128–31 Mosaic Systems, 172 multi-mode, 98 multistage countercurrent extraction, 31 mushrooms, 521 nanofiltration, 359–60, 393–4, 559, 610–12 nanofiltration membrane, 246–7 Natrix Separations, 172 nitro-keg product, 447 non-covalent imprinting, 150 norbixin, 298 Norit membrane, 443 nutraceutical industries electrodialytic phenomena and membrane technologies, 202–14 applications, 204–13 future trends, 213–14 principles, 203 nutraceuticals, 136, 148–9 novel adsorbents and approaches for separation, 148–9 membrane adsorbers and membrane chromatography, 169–72 mesoporous molecular sieves, 163–6 molecular imprinted polymers, 149–53 organic monoliths, 153–9 peptide affinity ligand and phage display methodology, 166–9 stimuli-responsive resins, 159–63 ohmic heating-assisted extractions, 83–6 apples and potatoes conductivity disintegration index Z, 84 juice expression and solute extraction, 84–6 principles, 83–4 oilseed processing extrusion preparation, 399–403 closed-wall extruders, 400–1 Dox-Hivex extruders, 402–3 dry extrusion, 402 extrusion before solvent extraction, 399–400
new developments, 403 origin of extrusion, 399 slotted-wall extruders, 401–2 mechanical pressing, 403–15 cages and shaft, 406 cone choke on screw press, 404 Desmet Ballestra press, 411 Dupps screw press, 411 Duvis press, 414 extrusion–screw press system, 410 French OM press, 412 full press suppliers, 408–13 Harburg-Freunberger press, 412 Insta-Pro extruder, 413 Insta-Pro screw press, 414 new developments in screw pressing, 405–8 other presses, 413–15 placing spacers, 406 screw press, 404 Victor 600 press, 409 oil from fruit pulps, 424–5 olive oil, 425 palm oil, 424–5 percolation solvent extraction, 415–22 extractor suppliers, 418–22 new developments in solvent extractors, 422 perforated belt extractor, 416–17, 420 perforated belt schematic, 421 rectangular loop extractor, 418 rectangular loop schematic, 419 rotary extractors, 416, 417 sliding cell extractor, 417–18 preparation, 397–9 cleaning, 397 cooking, 398–9 dehulling, 397 flaking, 398 separation technologies, 396–426 future trends, 425–6 solvent recovery, 422–4 effluent air, 423–4 effluent water, 424 marc, 423 miscella, 422–3 olive oil, 633 omega-3 fatty acids, 483–4 chromatographic methods, 486–8
© Woodhead Publishing Limited, 2010
Index 659
centrifugal partition chromatography, 488 concentration and purification, 483–502 enzymatic methods, 495–8 integrated methods, 498–501 low temperature fractional crystallisation, 488–90 microalgal biomass DHA, 499 supercritical-fluid extraction, 490–2 distillation methods, 492–5 crude fish oil purification, 494 molecular distillation, 493 urea adduction, 484–6 algal oil and its DHA concentrate, 485 ORAC test see oxygen radical absorbance capacity test organic monoliths, 153–9 advantages and disadvantages, 157 applications, 157–8 carbohydrates separation, 158 commercialisation, 158–9 functionalisation, 155–7 glycidyl methacrylate-co-ethylene glycol dimethacrylate monolith, 156 synthesis, 155 organophilic membranes, 229 organophilic pervaporation, 228, 236 osmotic distillation, 250–65, 391–3 applications, 261–5 concentrated fruit juices production, 263 heat transfer, 255–6 mass transfer aspects, 252–65 membranes and modules, 256–8, 259 typical membranes used, 259 operating conditions effect on the OD flux, 258, 260–1 process fundamentals, 250–2 concentration profile, 252 osmotic evaporation, 558, 560 see also osmotic distillation osmotic membrane distillation (OMD), 188 OTAP 92, 604–5 ovalbumin, 597–8, 604, 605 ovokinin, 602
ovomucin, 598–9 ovomucoid, 598 ovotransferrin, 373, 598, 604–5 oxygen radical absorbance capacity test, 212 panning, 168 passion fruit, 234 PDESol, 118 Peng–Robinson EOS, 16 peptide affinity ligands, 166–9 advantages and disadvantages, 167 commercialisation, 168–9 peptide identification, 167–8 peptides, 112, 363–6, 512, 515 percolation solvent extraction, 415–22 perforated belt extractor, 416–17, 420, 421 perlite, 439–40 permeate flux, 225 pervaporation aroma compounds recovery, 230–9 membranes selection, 227–9 methods sweeping-gas pervaporation, 222 thermopervaporation, 222 vacuum pervaporation, 222 principles, 220–1 typical experimental set-up, 221 principles and applications in food industry, 219–39 transport mechanism, 221–7 solute transport principle, 223 PFE, 61 see also pressurised fluid extraction phage display, 168 methodology, 166–9 phenolics, 511–14, 516–18, 535–6 phospholipids, 375 phosvitin, 600, 605 phosvitin phosphopeptides (PPP), 605 physically assisted extraction high-voltage electrical discharges, 86–90 solute extraction, 87–90 underwater HVED principles, 86–7 microwave-assisted extraction, 96–100 main applications, 98–100 principles, 96–8 reactors, 98
© Woodhead Publishing Limited, 2010
660 Index safety considerations, 100 ohmic heating-assisted extractions, 83–6 juice expression and solute extraction, 84–6 principles, 83–4 physical treatments combination, 100–2 principles and applications in food industry, 71–102 pulsed electric field-assisted extractions, 72–83 juice expression and solute extraction, 76–83 principles, 72–6 ultrasound-assisted extraction, 90–6 factors affecting UEA, 92–3 food ingredients, 95 hazard analysis critical control point, 95–6 main applications, 93–5 principles, 90–1 reactors, 92 plate-and-frame modules, 261 poly (N-isopropylacrylamide) (polyNIPAAm), 159 polydimethylsiloxane (PDMS), 229 polyether block polyamide (PEBA), 228 polyether sulphone membranes, 444–5 polyhalogenated persistant organic pollutants, 56 PolyHipe, 128–9 polymeric monoliths, 155, 157 polymers, 227 polymethyl octyl siloxane (POMS), 229 polyphenols, 525–6 polysaccharides, 563–4 polyvinyl polypyrrolidone (PVPP), 442 power ultrasound, 90, 91 Poynting correction, 13 PR EOS see Peng–Robinson EOS pre-cheeses, 353 prebiotics, 192 precipitation, 607–8 by ionic strength and/or pH modifications, 607 by organic solvents, 607–8 premature yeast flocculation (PYF), 435–6
Pressor, 409–10 pressure extractions, 71 pressurised fluid extraction applications, 56–9 1995 to 2008 publication record, 56 commercially available equipment, 61 design of your own equipment, 62–3 home-built dynamic PFE system, 62 environmental applications, 56–7 organometallic compounds, 57 pesticides, 56 pharmaceutical applications, 57 food and agricultural applications, 57–9 acylglycerols and sterols, 59 flavonoids, 57–8 other applications, 59 phenolic acids, 58 terpenoids, 58–9 tocopherols and tocotrienols, 59 future trends, 59–61 instrumentation and principles, 41–55 accuracy and precision, 54–5 basic instrumentation, 41–3 extraction curve calculation in red onion, 52 extraction pressure, 53 extraction solvents chemical properties, 47 extraction strategy, 43–7 extraction time, 52–3 household waste particle and sites where analytes might be found, 43 Irganox 1076 extraction, 45 olive oil liquid chromatogram, 55 selectivity during the extraction, 54, 55 solubility parameters for betulin, ethanol and water, 50 solvent selection, 47–50 static PFE system schematic diagram, 42 temperature effects, 50–2 principles and applications, 39–64 pressurised liquid extraction (PLE), 40, 538–47 pressurised solvent extraction (PSE), 40
© Woodhead Publishing Limited, 2010
Index 661 Primin process, 356 pro-oxidants, 508 probe system, 92 probiotics, 135 process chromatography, 118, 135 advances and applications in food, beverages and nutraceutical industries, 109–37 chromatography as a unit operation, 109–12 process-scale chromatography, 110 applications, 118–28 beers stabilisation, 122–3 glucose–fructose separation, 119–21 industrial scale chromatographic separations, 121 lactoferrin and lactoperoxidases operation, 126–7 lysozyme separation from egg white, 123–4 napin separation from rapeseed meal, 127–8 sucrose recovery from molasses, 121–2 whey protein from milk, 124–6 whey proteins typical concentration, 125 basic modelling, 117–18 mathematical modelling for chromatography, 119–20 basic principles, 113–18 chromatographic separation, 115 large-scale chromatographic fractionator operation, 118 modes of operation, 116–17 future trends, 135–7 high-value nutraceuticals, 135–6 process analytical techniques and control, 137 regulations, 136–7 list of abbreviations, 137–8 mode of operation, 132–5 continuous multi-column moving bed chromatography, 133 counter current chromatography, 132–4 expanded bed adsorption, 135 process control, 135 recent developments, 128–35
areas of innovation, 129 monoliths photographs, 130 novel ligands, 131–2 structured matrices and monoliths, 128–31 types of chromatography used, 113–16 basic interaction modes of chromatographic separation, 113 elution, 114 loading, 114 resin equilibration, 114 sanitisation, 115–16 sorbents requirements and main types, 116 washing, 114 Profi system, 444 protease hydrolysis, 558, 562 protein-A mimetic affinity peptides, 473 proteins, 512, 515 PSE, 61 pulse electric fields (PEF), 552, 556 pulsed electric field-assisted extractions, 72–83 juice expression and solute extraction, 76–83 apples, 81–2 grapes, 82 laboratory cell combining PEF treatment and juice expression, 77 other roots and tubers, 79–81 pilot belt press and juice yield and purity, 79–80 sugar beets, 76–8 yeast cells, 82–3 principles, 72–6 cell membrane electroporation mechanism, 73–4 characteristic time for red beet tissue half damage, 75 membrane electroporation, 73 PEF-treatment chambers, 76 plant tissue electrically induced damage, 74–6 purification methods acidic whey proteins, 454–62 a-lactalbumin, 461–2 adsorptive membrane separation, 458
© Woodhead Publishing Limited, 2010
662 Index
b-lactoglobulin, 454–60 bovine serum albumin, 462 chromatographic separations, 456–7 membrane separations, 455 other separation techniques, 459 basic whey proteins, 463–70 examples, 465–6 lactoferrin and lactoperoxidase, 464, 467–9 lysozyme, 469–70 dairy nutraceuticals, 450–74 future trends, 473–4 immunoglobulins, 471–3 affinity chromatography, 472–3 ion exchange chromatography, 471–2 membrane chromatography, 472 protein-A mimetic affinity peptides, 473 omega-3 fatty acids, 483–502 chromatographic methods, 486–8 distillation methods, 492–5 enzymatic methods, 495–8 integrated methods, 498–501 low temperature fractional crystallisation, 488–90 supercritical-fluid extraction, 490–2 urea adduction, 484–6 Q Sepharose Fast Flow, 608 RADHPF, 602 radical polymerisation, 155 reactive oxygen species (ROS), 521 rectangular loop extractor, 418, 419 Reflex rotary extractor, 418, 420 Reflex solvent extractor, 418 rennet coagulation, 354–5 resin purification, 565–7 resin screening, 111 reverse osmosis, 246, 247–8, 358–9, 393–4, 610–12 fruit juice industry, 185–7 reversed-phase high-performance liquid chromatography, 487, 610 Ricinus communis agglutinin (RCA), 131–2
Rosedown press see Sterling press rotary extractors, 416, 417 rubbery membranes, 228 S-ovalbumin, 598 saccharides concentration and fractionation, 191–5 separation set-up design, 192 operational modes, 194–5 feed solution concentration, 195 pH, 194 pressure, 195 temperature, 194 Sartorius membranes, 444 screw press, 403–5 seaweed, 520–1 secretory IgA, 470 selective elution method, 461 selective precipitation, 454–5 separation technologies, 341–69 brewing, 430–48 beer fining agents, 436–7 brewery products characteristics, 431–2 bulk beer filtration by membranes, 443–6 cleaning agents recovery, 446 dissolved gas control by membrane technology, 446–7 filter aid filtration and applications, 437–41 future trends, 447–8 regeneratable and reusable filter aids and applications, 441–2 technology and raw materials selection, 432 whirlpools and applications, 434–5 wort production in brewhouse, 433–4 yeast flocculation and applications, 435–6 dairy and egg processing, 341–76 dairy industry and compositions of dairy products, 343–7 egg products and composition, 369–71 egg white proteins extraction, 371–4 fractionation of individual proteins and peptides, 360–6
© Woodhead Publishing Limited, 2010
Index 663
future trends, 368–9, 375–6 milk, 347–51 standardisation and concentration of milk proteins, 351–4 treatment of effluents and technical fluids, 366–8 wheys and derivatives in cheese production, 357–60 whole casein isolation, 354–6 yolk components extraction, 374–5 fruit juice processing, 381–94 foods/fluids characteristics, 382–3 fruit juice concentrate, 386–94 product quality optimisation in product sector, 383–6 oilseed processing, 396–426 extrusion preparation, 399–403 future trends, 425–6 mechanical pressing, 403–15 oil from fruit pulps, 424–5 percolation solvent extraction, 415–22 preparation, 397–9 solvent recovery, 422–4 Sephadex G-200 column, 609 Sepharose 4B, 608 serum proteins, 344 fractionation, 361–3 microfiltration, 355–6 sheet filters, 441 short path distillation see molecular distillation silicalite-filled PDMS membrane, 236 silk fibroin, 564 simulated moving bed (SMB), 111, 132–4, 158, 467 single-mode, 98 single-stage extraction, 31 size-exclusion chromatography (SEC), 113, 469 stimuli-responsive resins applications, 162 sliding cell extractors, 417–18, 420, 422 slotted-wall extruders, 401–2 solubility, 634 solubility parameters, 44 solubility theory, 44 solute extraction, 85–6 solvent extraction, 71, 307–8, 526–31 solvent power, 24
sonication, 291 Soxhlet, 50, 52–3 apparatus, 95 extraction, 95, 102 Spherosil process, 358 squalene, 51 stack cookers, 398–9 static PFE, 41 Sterling press, 409 stimuli-responsive polymers, 159 stimuli-responsive resins, 159–63 applications, 160–3 ion-exchange chromatography, 160–1 commercialisation, 163 effect of critical temperature on poly-NIPAAm, 160 insulin chains and endorphin fragment mixture separation, 162 lactoferrin equilibrium adsorption isotherms, 161 synthesis, 159–60 Streamline SP, 609 subcritical-water extraction (SWE), 539–40 submerged membrane bioreactor, 317 Super Stripper System, 418 supercritical carbon dioxide, 621 supercritical carbon dioxide extraction (SCDE), 307–8 supercritical carbon dioxide fractionation, 362–3 supercritical fluid chromatography (SFC), 136 supercritical fluid extraction (SFE), 547–52, 556, 558 advantages, 4 commercial scale multistage system, 623 cycle processes for extraction, 21–6 extraction plant design, 26–30 diffusion, 29 energy consumption, 30 hydrodynamics, 29–30 mass transfer, 28–9 solids pre-treatment, 27–8 specific data, 26–7 thermodynamic data, 28 typical extraction curve, 29
© Woodhead Publishing Limited, 2010
664 Index
gas extraction process in temperature–entropy diagram solvent circuit in the compressor mode, 23 solvent circuit in the pump mode, 22 liquid-supercritical fluid equilibrium, 19–20 binary fluid systems phase behaviour classification, 19 high-pressure vapour–liquid equilibria thermodynamic modelling, 20 phase diagrams, 19–20 liquids extraction, 30–2 applications, 32 column under preparation for pressure test, 32 operation methods and apparatus, 31–2 lycopene from tomatoes, 619–40 co-solvent and modifiers on extraction, 631–4 factors affecting yield, 623–8 future trends, 639–40 pressure and temperature effects on antioxidant activity, 628–31 process, 622–3 solubility in supercritical fluids, 634–8 omega-3 fatty acid concentration and purification, 490–2 plant food antioxidants, 549–51 principles and applications in food industries, 3–36 extraction plant general flow sheet, 5 extractor closure, 7 extractors cascade for solid materials extraction, 7 high-pressure extractor solids during manufacturing, 6 high-pressure extractor unit, 6 multi-step separation, 8 number of compressed fluid extraction plants, 34 SFE processes maximum pressure, 34
specific processing coats vs operating pressure, 35 ultra high pressure extraction unit, 35 single-stage system with CO2, 622 solid-supercritical fluid equilibrium, 8–18 constants in cubic equations of state, 15 cubic equations of state, 14 phase diagrams, 8–11 pure solute fugacity calculation, 18 simplest solid–supercritical fluid equilibrium, 9 solid–liquid–gas equilibrium lines, 10 solid–liquid–gas phase equilibria, 11 solid–liquid–gas phase lines, 12 thermodynamic modelling, 11–17 solids extraction, 26–30 application, 30 solute separation in extraction process, 24–6 expansion, 25 mass separating agent, 25–6 reduced solvent power, 24–5 solvent cycle, 21–4 compressor process, 23–4 extraction process using supercritical fluid, 21 pump process, 22–3 thermodynamics fundamentals, 8–20 Superose 6, 608 Superose 12 HR 10/30 column, 608 SuperProDesign, 118 suspension polymerisation methods, 152 sweeping gas membrane distillation (SGMD), 188, 266 sweeping-gas pervaporation, 222 tannins, 512 tartaric acid, 206 temperature polarisation, 227, 264 Termamyl, 327–8 terpenes, 512 tetrahydrofuran, 51 TF200, 257 TF450, 257
© Woodhead Publishing Limited, 2010
Index 665 TF1000, 257 thermal evaporation, 245 thermo-electroplasmolysis, 84 thermopervaporation, 222 tocopherols, 524 toluene, 638 tomatoes supercritical-fluid extraction of lycopene, 619–40 effects of co-solvent and modifiers, 631–4 effects of pressure and temperature, 628–31 factors affecting extraction yield, 623–8 future trends, 639–40 lycopene solubility, 634–8 process, 622–3 tricin, 564–5 Trolox, 515 tunnel pasteurisation, 431 twin-screw press, 425 ultrafiltration, 182, 351–4, 460, 472, 557, 558, 559, 561, 610–12 ultrafiltration membranes, 246 ultrasonic cleaning bath, 92 ultrasound and MW-assisted extraction (UMAE), 99–100 ultrasound-assisted extraction (UAE), 90–6, 552 factors affecting UAE, 92–3 ultrasound-induced cell damage mechanism, 93 hazard analysis critical control point, 95–6 good maintenance plan, 96 laboratory, industrial and ultrasound extraction reactors, 92 main applications, 93–5 antioxidants, 94–5 essential oils and aromas, 93–4 food ingredients, 95 oil and fat, 95 orange peels rapid sono-extraction in alcoholic beverages, 94 principles, 90–1
ultrasonic cavitation phenomena, 91 reactors, 92 ultrasound-assisted maceration, 94 uniform transmembrane pressure (UTP), 349 urea adduction, 484–6 vacuum evaporation, 185, 187 vacuum membrane distillation (VMD), 188, 266, 273 vacuum pervaporation, 222 value-added food products separation by colloidal gas aphrons flotation, 284–309 CGA applications, 293–307 CGA properties and applications, 285–93 future trends, 308–9 industrial application feasibility, 307–8 van der Waals one fluid mixing rules, 15–16 van der Waals one fluid model, 17 vegetable oils, 638 vibrating membrane module, 196 Victor 600 press, 408, 409 Vistec process, 358 vitamin E, 512, 524, 525 vitellogenin II, 600 water, 632–3 whey, 346–7, 357–60 separation techniques, 357–60 concentration and demineralisation, 358–60 demineralisation, 359 serum proteins concentration, 357–8 sweet and acid wheys composition, 347 whey proteins, 451–70 whirlpools, 434–5 Wong and Sandler mixing rules, 17 wort, 432–4 yeast flocculation, 435–6 yellow passion fruit, 212
© Woodhead Publishing Limited, 2010
