NEW INSIGHTS INTO MEMBRANE SCIENCE AND TECHNOLOGY: POLYMERIC AND BIOFUNCTIONAL MEMBRANES
Membrane Science and Technology Series Volume 1 : Pervaporation Membrane Separation Processes Edited by R.Y.M. Huang (1991) Volume 2 : Membrane Separations Technology, Principles and Applications Edited by R.D. Noble and S.A. Stern (1995) Volume 3 : Inorganc Membranes for Separation and Reaction By H.R Hsieh (1996) Volume 4 : Fundamentals of Inorganic Membrane Science and Technology Edited by A.J. Burggraaf and L. Cot (1996) Volume 5 : Membrane Biophysics Edited by H. Ti Tien and A. Ottova-Leitmannova (2000) Volume 6 : Recent Advances in Gas Separation by Microporous Ceramic Membranes Edited by N.K. Kanellopoulos (2000) Volume 7 : Planar Lipid Bilayers (BLMs) and their Applications Edited by H.T. Tien and A. Ottova-Leitmannova (2003) Volume 8 : New Insights into Membrane Science and Technology: Polymeric and Biofunctional Membranes Edited by D. Bhattacharyya and D.A. Butterfield
Membrane Science and Technology Series, 8
NEW INSIGHTS INTO MEMBRANE SCIENCE AND TECHNOLOGY: POLYMERIC AND BIOFUNCTIONAL M E M B I ~ N E S Edited
by
Dibakar Bhattacharyya Department of Chemical and Materials Engineering and Center of Membrane Sciences University of Kentucky, Lexington, KY 40506-0046, USA and
D. Allan Butterfield Department of Chemistry and Center of Membrane Sciences University of Kentucky, Lexington, KY 40506-0059, USA
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TO:
Bhattacharyya: Gloria and Anita, my wife and daughter, for their support and understanding," my graduate students, for making teaching and research life to be very exciting. Butterfield: Marcia and Nyasha, my wife and daughter, who have been so inspirational, encouraging, and loving.
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Preface Membrane techniques provide a broad science and technology base with applications ranging from water purification to chemical/biomolecule synthesis, material recovery, medical devices, to nano-domain interaction- based sensors and highly selective separations. Although there are several books in the traditional membrane field, there is a great need for a highly comprehensive book containing advances in both synthetic and biofunctional/bimimetic membranes including non-invasive characterization to biomedical devices. Over the years the membrane field has advanced from the development of thinfilm composite membranes for desalination to recognition-based separation and reactions by taking advantage of biomolecular interactions in the nano-domain. This refereed book covers materials from highly respected researchers with topics ranging from membrane transport models to non-invasive characterization, functionalized material, biomedical devices, to sensors and environmental applications. Nineteen chapters in this special book are invited, refereed papers, and mostly based on the 2001 North American Membrane Society Annual Meeting held in Lexington, KY, organized and hosted by the University of Kentucky Center of Membrane Sciences. The book is divided into two sections. Section I, is subdivided into three areas, Advances in Membrane Transport~ouling, Imaging Techniques, and Contacting Devices. Section II deals with Functional Membranes and Materials for Biocatalysis, Separation, and Analysis and is further subdivided into three areas ranging from biofunctional membranes to sensors. Authors and co-authors are from various fields including chemistry, chemical engineering, mechanical engineering, biomedical engineering, biotechnology, and environmental engineering. This book is highly multidisciplinary in nature and should be highly valuable to scientists and engineers involved in activities ranging from separations/reactions, to advanced biofunctional materials, to contactor designs, to the general field of membrane science and technology. Students and faculty members around the world should find this to be an excellent reference book for courses ranging from traditional separation, to bioscience/engineering, to formal membrane courses. Each chapter of this book was peer-reviewed, and we would like to give special thanks to the reviewers for this book, including various authors and co-authors of the chapters, and Drs. Kloos (Osmonics Corporation) and Hestekin (Argonne National Lab). Thanks also go to the NAMS Board of Directors, and in particular to NAMS presidents Drs. Pushpinder Puri and Glenn Lipscomb for their encouragement of the book. Mollie Fraim of the University of Kentucky °. Vll
Center of Membrane Sciences did an extraordinary amount of work with us to bring this book to fruition, and we would like to express our sincere appreciation and thanks to Ms. Fraim. The editors would also like to thank Mike Lundin (chemical engineering student) for his hard work in putting the book in right format.
Dibakar Bhattacharyya, Ph.D. Department of Chemical and Materials Engineering and Center of Membrane Sciences University of Kentucky Lexington, KY D. Allan Butteriield, Ph.D. Department of Chemistry and Center of Membrane Sciences University of Kentucky Lexington, KY
viii
About the Editors Dibakar Bhattacharyya (DB) is the University of Kentucky Alumni Professor of Chemical Engineering and a Fellow of the American Institute of Chemical Engineers. He obtained his B.S. (Jadavpur University) and M.S. ( Northwestern University, Evanston, IL) in Chemical Engineering, and his Ph.D. in Environmental Engineering from the Illinois Institute of Technology. He is the Co-Founder of the Center for Membrane Sciences at the University of Kentucky. He and his students have advanced passive membrane applications to functionalized poly-ligand membranes for high capacity metal capture to site-specific biocatalysis for organic degradation to membrane-based nanoparticle synthesis for dechlorination reactions. He has published over 140 refereed journal articles and book chapters, and has recently received three U.S. Patents (two on Functionalized Membranes for Toxic Metal Capture, and one on hazardous waste destruction technology). Dr. Bhattacharyya has mentored many graduate and undergraduate students in the area of water and wastewater related research and membrane separation. For his research, Dr. Bhattacharyya has received funding from U.S. EPA, DoD, NSF, NIEHS-SBRP, Dow Chemical, Glaxo SmithKline, Eastman Chemical Co., Daramic, Inc., etc. He has received a number of awards for his research and educational accomplishments, including the Larry K. Cecil AIChE Environmental Division Award, the Kentucky Academy of Sciences Distinguished Scientist Award, Henry M. Lutes Award for Outstanding Undergraduate Engineering Educator, AIChE Outstanding Student Chapter Counselor Awards, and the University of Kentucky Great Teacher (1984 and 1996) Awards. D. Allan Butterfield received his B.A. in Chemistry from the University of Maine in 1968. Following three years of teaching Africans chemistry and mathematics in Zimbabwe, he entered Duke University, receiving the Ph.D. in Physical Chemistry in 1974. A NIH Postdoctoral Fellowship at the Duke University Neuroscience Program followed. Dr. Butterfield joined the Chemistry Department at the University of Kentucky soon after and was quickly promoted to Full Professor in 1983. In 1986, he and several others, but principally Professor Dibakar Bhattacharyya, formed the University of Kentucky Center of Membrane Sciences, and Professor Butterfield has been Director from its inception. Continuous federal funding from NIH, NSF, and DoD has supported his research on membrane structure and function in neurodegenerative disorders and on applications of biofunctional membranes to important societal problems. He and his students have published over 260 refereed papers. Professor Butterfield received the Southern Chemist Award from the American Chemical Society. He has directed the graduate careers of over 50 Ph.D. and M.S. students, and in 1998 Professor Butterfield received the Presidential Award for Excellence for Science, Mathematics, and Engineering Mentoring from President Clinton in the White House. This award was his efforts to increase the number of female and Appalachian Ph.D. students in Chemistry, both groups being highly under represented in this discipline. In 2002, the University of Kentucky Board of Trustees appointed Dr. Butterfield the Alumni Professor of Chemistry.
ix
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Contents Preface ................................................................................................
vii
About the Editors . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .
ix
List of Contributors . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .
xv
A d v a n c e s in M e m b r a n e T r a n s p o r t / F o u l i n g , I m a g i n g T e c h n i q u e s , a n d C o n t a c t i n g D e v i c e s .............................................................................. 1
Membrane Transport Models, Fouling, and Formation ..................... 3
Mass Transfer in Axial Flows through Randomly Packed Fiber Bundles ......................................................................... 5 L. Bao, G.G. Lipscomb Fouling Phenomena during Microfiltration: Effects of Pore Blockage, Cake Filtration, and Membrane Morphology ....... 27 A.L. Zydney, C. Ho, W. Yuan Differential Scanning Calorimetry and rheological experiments to study membrane formation via thermallyinduced phase-separation .................................................................. 45 P.C. van der Heijden, M.H.V. Mulder, M. Wessling Non-lnvasive Characterizations o f Membrane Fluid Transport and Fouling ........................................................................................ 63
Study of Membrane Fouling and Cleaning in Spiral Wound Modules Using Ultrasonic TimeDomain Reflectometry ...................................................................... 65 Zh.-X. Zhang, A.R. Greenberg, W.B. Krantz, G.-Y. Chai Nonintrusive Characterization of Fluid Transport Phenomena in Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach ......................... 89 C.K. Poh, P.A. Hardy, W.R. Clark, D. Gao Membrane Contactors and Environmental Applications ................ 123
Industrial applications and opportunities for membrane contactors ....................................................................................... 125 R. Klaassen, P.H.M. Feron, R. van der Vaart, A.E. Jansen xi
Membrane Contactors: Recent Developments ................................ 147 A.S. Kovvali, K.K. Sirkar
Membrane Aromatic Recovery System (MARS)A New Process for Recovering Phenols and Aromatic Amines from Aqueous Streams ....................................... 165 F.C. Ferreira, A. Livingston, S. Han, A. Boam, S. Zhang Functional Membranes and Materials for Biocatalysis, Separation, and Analysis ............................................................................................ 183
Biofunctional Membranes and Biomedical Devices ........................ 185 Membrane Bioreactors for Biotechnology and Medical Applications ................................................................... 187 L. Giomo, L. De Bartolo, E. Drioli Structural and Performance Characteristics of Hemodialysis Membranes ............................................................ 219 D. Gao, W.R. Clark Biofunctional Membranes: Site-Specifically Immobilized Enzyme Arrays ........................................................ 233 D.A. Buttertield, D. Bhattacharyya Biocatalytic Membrane Reactor with Continuous Removal of Organic Acids by Electrodialysis .............................. 241 H.C. Ferraz, T.L.M. Alves, C.P. Borges Use of Micro-Porous Affinity Membranes for Protein Purification: A Case Study ............................................... 263 F. Cattoli, G.C. Sarti Economic Production of Biopharmaceuticals by High-Speed Membrane Adsorbers ............................................... 283 W. Demmer, S. Fischer-Frtihholz, D. NuBbaumer, D. Melzner Functionalized Membranes f o r Separations and Reactions ............ 297 Polymer Grafted Membranes ....................................................... 299 S.M.C. Ritchie
xii
Functionalized Membranes for Tunable Separations and Toxic Metal Capture ............................................................... 329 A.M. Hollman, D. Bhattacharyya The Design of High Performance, Gel-Filled Nanofiltration Membranes ........................................................... 353 R.F. Childs, A.M. Mika Sensors ............................................................................................. 377 Membranes for the Development of Biosensors ........................... 379 V.G. Gavalas, J. Wang, L.G. Bachas
Ion-Partitioning Membranes as Electroactive Elements for the Development of a Novel Cation-Selective CHEMFET Sensor System ................................ 393 E.A. Moschou, N.A. Chaniotakis
Index ........................................................................................................ 415
xiii
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List of Contributors
T.L.M. Alves Chemical Engineering Program- COPPE Federal University of Rio de Janeiro, P.O. Box: 68502 21945-970, Rio de Janeiro R J - Brazil Leonidas G. Bachas Department of Chemistry and Center of Membrane Sciences University of Kentucky Lexington, Kentucky 40506-0055 USA E-mail: bachas@pop, uky. edu Lihong Bao Chemical & Environmental Engineering Department Mail Stop 305 The University of Toledo Toledo, Ohio 43606-3390 USA Dibakar Bhattacharyya Department of Chemical and Materials Engineering and Center of Membrane Sciences University of Kentucky Lexington, KY 40506-0046 USA E-mail: db@engr, uky. edu Andrew Boam Membrane Extraction Technology Ltd Imperial College London SW7 2BY UK C.P. Borges Chemical Engineering Program - COPPE Federal University of Rio de Janeiro, P.O.Box: 68502 21945-970, Rio de Janeiro R J - Brazil E-mail: cristiano@peq,coppe, ufrj.br
XV
D. Allan Butterfield Department of Chemistry and Center of Membrane Sciences University of Kentucky Lexington, KY 40506-0059 USA E-mail: dabcns@uky, edu Francesca Cattoli DICMA - Dipartimento di Ingegneria Chimica, Mineraria e delle Tecnologie Ambientali Alma Mater Studiorum - Universit~t di Bologna, viale Risorgimento 2 40136 Bologna Italy G.-Y. Chai Department of Chemical Engineering University of Colorado Boulder, Boulder CO 80309-0424 USA Nikolas A. Chaniotakis Laboratory of Analytical Chemistry Department of Chemistry University of Crete, 71409 Iraklion, Crete Greece Ronald F. Childs Department of Chemistry McMaster University Hamilton, ON, L8S 4M1 Canada E-mail: rchilds@mcmaster, ca William R. Clark, M.D. Renal Division Baxter Healthcare Corporation McGaw Park, IL Nephrology Division Indiana University School of Medicine Indianapolis, IN USA E-mail: bill_clar~axter, com
xvi
L. De Bartolo Research Institute on Membranes and Modelling of Chemical Reactors, IRMERC-CNR c/o University of Calabria via P. Bucci cubo 17/C, 87030 Rende (CS) Italy W. Demmer Sartorius AG G/Sttingen Germany E. Drioli Research Institute on Membranes and Modelling of Chemical Reactors IRMERC-CNR c/o University of Calabria via P. Bucci cubo 17/C, 87030 Rende (CS) Italy E-mail: e. drioli@irmerc, cs. cnr. it
P.H.M. Feron TNO Institute of Environmental Sciences Energy Research and Process Innovation Department Chemical Engineering PO Box 342 7300 AH Apeldoorn The Netherlands H.C. Ferraz Chemical Engineering Program - COPPE Federal University of Rio de Janeiro, P.O.Box: 68502 21945-970, Rio de Janeiro R J - Brazil S. Fischer-Frtihholz Sartorius AG Gfttingen Germany Frederico Castelo Ferreira Department of Chemical Engineering and Chemical Technology Imperial College of Science Technology and Medicine Prince Consort Road London SW7 2BY UK
xvii
Dayong Gao Department of Mechanical Engineering Center for Biomedical Engineering University of Kentucky Lexington, KY USA E-mail: dgao@engr, uky. edu Vasilis G. Gavalas Department of Chemistry University of Kentucky Lexington, Kentucky 40506-0055 USA L. Giomo Research Institute on Membranes and Modelling of Chemical Reactors, IRMERC-CNR c/o University of Calabria via P. Bucci cubo 17/C, 87030 Rende (CS) Italy A.R. Greenberg Department of Mechanical Engineering University of Colorado Boulder, Boulder CO 80309-0427 USA E-mail: greenbea@,spot, colorado, edu Shejiao Han Department of Chemical Engineering and Chemical Technology Imperial College of Science Technology and Medicine Prince Consort Road London SW7 2BY UK Peter A. Hardy Center for Biomedical Engineering University of Kentucky Lexington, KY USA Chia-Chi Ho Department of Chemical Engineering University of Cincinnati Cincinnati, OH 45221 USA
xviii
Aaron M. Hollman Department of Chemical & Materials Engineering University of Kentucky Lexington, Kentucky 40506-0046 USA A.E. Jansen TNO Institute of Environmental Sciences Energy Research and Process Innovation Department Chemical Engineering PO Box 342 7300 AH Apeldoom The Netherlands R. Klaassen TNO Institute of Environmental Sciences Energy Research and Process Innovation Department Chemical Engineering PO Box 342 7300 AH Apeldoorn The Netherlands E-mail: r. klaassen@mep, tno.nl A. Sarma Kovvali Department of Chemical Engineering New Jersey Institute of Technology Newark, NJ 07102 USA W.B. Krantz Department of Chemical Engineering University of Cincinnati Cincinnati, Ohio 45221-0171 USA G. Glenn Lipscomb Chemical & Environmental Engineering Department Mail Stop 305 The University of Toledo Toledo, Ohio 43606-3390 E-mail:
[email protected]
xix
Andrew Livingston Membrane Extraction Technology Ltd. Imperial College London SW7 2BY UK and
Department of Chemical Engineering and Chemical Technology Imperial College of Science, Technology and Medicine Prince Consort Road London SW7 2BY UK E-mail: a. livingston@ic, ac. uk
D. Melzner Sartorius AG Gfttingen Germany E-mail: dieter.melzner@sartorius, com
Alicja M. Mika Department of Chemistry McMaster University Hamilton, ON, L8S 4M1 Canada Elizabeth A. Moschou Laboratory of Analytical Chemistry Department of Chemistry University of Crete, 71409 Iraklion Greece Current address: Department of Chemistry University of Kentucky Lexington KY 40506-0055 USA E-mail: l_moschou@yahoo, com
M.H.V. Mulder University of Twente PO Box 217 7500 AE Enschede The Netherlands D. NuBbaumer Sartorius AG Grttingen Germany
XX
Chum K. Poh Department of Mechanical Engineering University of Kentucky Lexington, KY USA S.M.C. Ritchie Department of Chemical Engineering University of Alabama Tuscaloosa, AL, 35487-0203 USA E-mail: sritchie@bama, ua. edu Giulio C. Sarti DICMA - Dipartimento di Ingegneria Chimica Mineraria e delle Tecnologie Ambientali Alma Mater Studiorum - Universith di Bologna viale Risorgimento 2 40136 Bologna Italy E-mail: giulio.sarti@,mail, ing. unibo, it Kamalesh K. Sirkar Department of Chemical Engineering New Jersey Institute of Technology Newark, NJ 07102 USA E-mail: sirkar@adm, nj it. edu P.C. van der Heijden University of Twente PO Box 217 7500 AE Enschede The Netherlands E-mail: pcvdheijden@hetnet, nl R. van der Vaart TNO Institute of Environmental Sciences Energy Research and Process Innovation Department Chemical Engineering PO Box 342 7300 AH Apeldoom The Netherlands Jianquan Wang Department of Chemistry University of Kentucky Lexington, Kentucky 40506-0055 USA xxi
M. Wessling University of Twente PO Box 217 7500 AE Enschede The Netherlands Wei Yuan Innovasep Technologies 420 Maple Street Marlborough, MA 01752 USA Shengfu Zhang Membrane Extraction Technology Ltd Imperial College London SW7 2BY UK Zh.-X. Zhang Department of Chemical Engineering University of Colorado Boulder, Boulder CO 80309-0424 USA Andrew L. Zydney Department of Chemical Engineering The Pennsylvania State University University Park, PA 16802 USA E-mail:
[email protected], edu
xxii
Advances in Membrane Transport/Fouling, Imaging Techniques, and Contacting Devices
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Membrane Transport Models, Fouling, and Formation
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New Insights into Membrane Science and Technology: Polymeric and Biofunctional Membranes D. Bhattacharyya and D.A. Butterfield (Editors) © 2003 Elsevier Science B.V. All rights reserved.
Chapter 1
Mass transfer in axial flows through randomly packed fiber bundles L. Bao and G, G. Lipscomb* Chemical & Environmental Engineering Department, Mail Stop 305, The University of Toledo, Toledo, Ohio 43606-3390 *Corresponding Author, e-mail:
[email protected], Phone: (419) 530-8088, Fax: (419) 530-8086 1.
ABSTRACT
The literature contains numerous correlations of mass transfer coefficients for axial flows through randomly packed fiber bundles. The predictions of these correlations can differ by an order of magnitude. This large variation severely limits the usefulness of the correlations for design purposes and confounds an understanding of the underlying physics. While the literature suggests randomness in fiber packing may contribute to the variation, a rigorous analysis of mass transfer in randomly packed bundles has not been reported. The results of such analyses are summarized here for uniform wall concentration and uniform wall mass flux boundary conditions. The results indicate that channeling through randomly packed bundles can dramatically reduce mass transfer coefficients relative to regularly packed bundles, especially m the welldeveloped limit. However, experimental mass transfer coefficients are significantly different from predicted values for regularly and randomly packed bundles; the literature contains correlations that predict much higher and much lower values. This suggests that other factors can control performance such as cross-flow regions that exist near shell ports or poor fluid distribution from the shell inlet port. 2.
INTRODUCTION
Hollow fiber membrane modules are used for a wide range of applications including dialysis [1], filtration [2], eontaetors [3], and gas separations [4]. Typically, the modules consist of a bundle of randomly packed hollow fibers. The ends of the fiber bundle are potted to form tubesheets - slicing the tubesheets provides access to the fiber interior or lumens. The bundle is inserted into a case and a seal formed between the tubesheet and ease. Ports on the
Mars TranafcrIn Axial Flm Tlmugb Rmdaoly Packed Fiber B d e d - Lipscomb
periphery of the case permit fluid introduction and removal from the shell space (the region outside the fibers) while headers on either end permit fluid introduction and removal from the lumen space (the region inside the fibers). The tubesheet prevents the direct mixing of the shell and lumen fluids. Figure 1 illustrates a typical module. In operation, mass transfer occurs between the fluids flowing through the lumen and shell and the module functions llke the mass exchange equivalent of a shell-and-tube heat exchanger. The material used to produce the hollow fiber membranes can facilitate the desired separation - ideally, the membrane would selectively permeate one or more components of the feed streams at high rates. Through proper choice of membrane material and operating conditions, one can separate the components of a feed stream into hgh-purity product streams. Unfortunately, the efficiency and throughput of these devices can be limited not by the membrane material but by concentration boundary layers that develop in the lumen and shell spaces. The literature [ 5 ] indlcates that the effect
Figure 1. Schematic of a typical hollow fiber membrane module and the case that holds it.
of lumen boundary layers is described well by the Graetz solution [6] for mass transfer coefficients in pipe flows. However, there is much greater uncertainty about the effect of shell boundary layers. Theoretical analyses and experimental measurements of shell-side mass transfer coefficients have been confounded by the complex geometry of the shell and the flows within it. This complexity arises primarily from four factors: 1) fibers are packed randomly in the module whch leads to irregular flow domains that can be described only statistically (i.e., one can speclfy a packing fraction but not the exact location of the fibers) in contrast to the single, well-defined domain that exists in regularly packed bundles; 2) the fibers may not lie parallel to each other which can lead to a transverse flow (cross-flow) component and a velocity field with three components that varies in all three coordinate duections in contrast to a uniaxial velocity field that varies in two coordinate directions in regularly packed bundles; 3) the need to introduce and remove fluid from the shell leads to cross-flow regions and axial-flow regions - concentration boundary layer growth depends strongly on the relative size of these regions but most analyses of heat and mass transfer assume only axial flow; and 4) fluid
6
Mass Transfer In Axial Flows Through Randomly Packed Fiber Bundles - Lipscomb
distribution into the bundle may not be uniform if the entrance and exit ports are not well designed- locally flows may be higher or lower than expected based on the assumption of a uniform axial pressure gradient in the bundle (This has been seen experimentally in hemodialyzers [7] where flows appear higher in that part of the bundle closest to the inlet port.). The last two factors can affect performance in both randomly and regularly packed fiber bundles. All four are illustrated in Figure 2. a) cross-sections of bundles with random (left) and regular square (fight) fiber packings o• -N • • o•i O %
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Figure 2. Four complexities associated with shell flow through hollow fiber membrane modules: a) random fiber packing, b) transverse flows arising from non-parallel fibers, c) cross-flow regions near shell inlet port, and d) poor distribution from shell inlet port.
Mass Transfer In Axial Flows Through Randomly Packed Fiber Bundles -
Lipscomb
Despite these challenges, the literature offers a number of mass transfer coefficient correlations nominally for axial flows through randomly packed hollow fiber bundles [8-13]. Lipnizki and Field [14] provide an excellent summary and discussion of the reported expressions. The proposed correlations are summarized in Table 1 in terms of the mean Sherwood number: Shr,,=2Rkm/D where R is the fiber radius, km the mean mass transfer coefficient and D the diffusion coefficient. The Sherwood number is a function of the Reynolds number (Re=2RVb,O/lt), Schmidt number (Sc=ldp/D), fiber length (L), and fiber packing fraction (¢~= total fiber cross-sectional area / module cross-sectional area). Here, V5 is the bulk fluid velocity in the shell, ,o the fluid density, and/t the fluid viscosity. Table 1. Literature mass transfer coefficient correlations for axial flows through hollow fiber bundles and the reported range of conditions used to obtain the correlations. In all cases the Schmidt number is approximately 500-1000. Note [15] that the correlation for Prasad and Sirkar [9] has been modified to provide a better fit to their data. Additionally, (1-¢)/~ is the ratio of the hydraulic diameter to the fiber diameter. Correlation
[8]
/, ~'xo.s6/,,,,,~o.93 Sh,,, = 1.25~2-~ l ~ ~ Re°93Sc °'33
[9]
Shin = 11.5
ll-',/[
(1_ +t_._~. j
[10] Shr,, = (0.53-0.58~tL~) [11] Sh,,, = 8.71
ReO.6ScO.33
0001
0
0001
,('1- ~'~-°'2-°16~ Re°'S-°'16~Sc °'33
Sh,,, = (0.31¢2 -0.34~+0.10
Re°'9S¢°'33
0.003
0.004 30
Re°'74Sc °'33
2R/L 0.0020.006
0 00
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[12] Sh. : 0.090-¢~-7J [13]
100~ Re(1-~)/ # 3-26 0.5-500
0-0.2
0.002
35-80 0-0.5
0.0030.005
10-70 -~0-200
-~0.006
Figure 3 compares the predictions of the mean Sherwood number as a function of Graetz number (defined here as Gz = Re*Sc*(rcR/2L)*(1/~-1) to facilitate comparison with the theory described in the next section) for correlations valid over the range of conditions shown or that require limited extrapolation. The predictions for a regular square packing (constant wall mass flux boundary condition) are given as well; values for triangular packing are just
Lipscomb
Mass Transfer In Axial Flows Through Randomly Packcxl Fiber Bundles -
slightly higher. Clearly, there is significant variation between the correlations. Additionally, the correlations predict values that are b o t h much high and much lower than that for a regular packing. The literature attributes the large differences in mass transfer coefficients to the above four factors. However, a rigorous analysis of the effect of each is not available, and only the first (the effect of random fiber packing) has received much attention. Rogers and Long [16], Wu and Chen [13], and Lipnizki and Field [14] describe similar ad hoe approaches to predicting the effect of random fiber packing. Each set of authors introduces randomness by surrounding a representative set of fibers with an annular region of varying cross-sectional area; the first two pairs of authors use Voronoi tessellation to generate polygonal regions while the last assumes concentric circular regions, see Figure 4. The 10 2
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. 10 3
Gz
Figure 3. Comparison of literature correlations for mean shell-side mass transfer coefficients in randomly packed bundles with theoretical predictions for regular square packing for []=0.3, Sc=1000, and 2R/L=0.003. The sources ofthe correlations are: o Yang and Cussler [8], + Prasad and Sirkar [9], * Costello et al. [10], x Viegas et al. [11], 0 Gawronski and Wrzesinska [12], and small dot Wu and Chen [13]. The [] represent theoretical predictions for a square packing with uniform wall mass flux from Miyatake and Iwashita [23]. flow through each region is calculated from an appropriate friction factorReynolds number relationship using the region's hydraulic diameter and assuming the pressure drop across each region is the same. Each pair of authors predict a reduction in mass transfer coefficient due to randomness, but the effect of their ad hoe approach on the results is unclear. Results from rigorous analyses of the effect of random packing on mass transfer coefficients for axial flows through fiber bundles are summarized here. As found previously, randomness reduces mass transfer coefficients due to
Mass Transfer In Axial Flows Through Randomly Packed Fiber Bundles - Lip~omb
channeling in the fiber bundle. The results provide a baseline for evaluation of experimental results and suggest most experimental correlations do not reflect mass transfer in axial flows - either the shell-side flow possessed a significant cross-flow component or flow distribution from the shell ports into the bundle was poor. b) Surrounding each fiber with concentric circular regions
a) Using Voronoi tessellation to surround individual fibers with polygons
Figure 4. Approximate approaches for evaluating the effect of random packing on module performance - mass transfer rates are calculated for each fiber by using the hydraulic diameter of the surrounding shell region to calculate a flow rate and mass transfer coefficient for that region.
3. THEORY Commercial modules contain from 10,000 to over 1 million fibers. Unfortunately, finite computational resources limit the ability to calculate velocity and concentration fields within such large bundles. To reduce computational complexity, we use a unit cell containing a finite number of fibers to represent an infinite, spatially periodic bundle. This unit cell is translated in each co-ordinate direction to generate the bundle; the fibers lie parallel to the zaxis and the unit cell is translated along the x- and y-axes as illustrated in Figure 5. The value of a field variable at a point (x, y) in the unit cell is the same as the value at a point (x + ilx, y + fly) in the infinite bundle where i a n d j are integers and lx and ly represent the unit vectors in the x and y co-ordinate directions, respectively. Fiber packing fraction and type of packing (i.e., a regular or random packing) determine unit cell dimensions. Fiber centers are placed on a regular grid to produce regular packings while fibers centers are placed in a random, sequential fashion to produce random packings. The size and properties of all fibers are assumed to be identical. The random media produced by this method are representative of infinite random media only if the unit cell contains a sufficient number of fibers. 10
Mass Transfer In Axial Flows Through Randomly Packed Fiber Bundles - Lipscomb
Calculations with increasing fiber numbers are used to verify that the calculated results do not depend on fiber number- up to 100 fibers were used in the work
lb-
Figure 5. Use of a unit cell to represent an infinite, randomly packed fiber bundle. The unit cell is indicated by the bold square and contains nine solid fibers. The solid lines emanating from the unit cell indicate the required translation to produce an infinite medium. The numbers 1 and 2 correspond to a pair of periodic boundaries: a denotes the top and bottom while b denotes the left and right.
reported here. However, given sufficient computational resources, the algorithms could be used for arbitrarily large fiber numbers. Additionally, one may use this approach to generate bundles containing non-uniform fibers if the unit cell contains a sufficiently large number of fibers to accurately represent the statistical distribution. This would further increase the required computational resources. The governing conservation of momentum and mass equations are solved given the following assumptions: 1. Fluid flow through the fiber bundle is axial, laminar, isothermal and welldeveloped; velocity field entrance lengths are more than an order of magnitude smaller than concentration field entrance lengths [14]. 2. Mass transfer rates are sufficiently low that shell flow rates are constant along the fiber bundle. 3. The wall concentration or wall mass flux along the surface of each fiber in the mass transfer region is constant. 4. Mass transfer by axial diffusion is negligible relative to convective mass transfer. 5. Diffusion is Fickian. 44
Mass Transfer In Axial Flows Through Randomly Packed Fiber Bundles -
Lipscomb
6. Physical properties of the fluid are constant. These assumptions are equivalent to those used to obtain the Graetz solution for mass transfer in pipes and allow one to write the conservation of momentum and mass equations as"
02V 02V
--+
ax 7
=1
(1)
0¢ 02¢ 02¢
v-- = --+--
Oz
Ox2
(2)
03,2
where v is the dimensionless axial velocity, c the dimensionless concentration, and x, y, and z the dimensionless coordinates. The dimensionless variables are defined as
V V= ~R2 , 0p
X
Y
x=-~-,
y=-~,
/t0Z
ZD Z=R40p
(3 /
p0Z
where P is the pressure. For a constant wall concentration boundary condition, c is given b y C -C w Co - C w
(4a)
where Co is the inlet fluid concentration and Cw the specified wall concentration. For a constant wall flux boundary condition, c is given by: C -C O c=~
(4b)
J~R/D
where Jw is the specified wall mass flux. The boundary conditions for Eq. (1) and (2) are: c-1 v = 0,
or
c=0
c = 0,
(5)
z=0, f o r a l l x , y or
ac/an = 1 along fiber surface
19
(6)
Mass Transfer In Axial Flows Through Randomly Packed Fiber Bundles - Lipscomb
=q=, cl, =cl=
(7a) (7b)
1
2
1
2
where n is the unit outward normal direction along the boundary of the solution domain. The first of the two values for the concentration boundary condition given in (5) and (6) is for the constant wall concentration case while the second is for the constant wall mass flux case. The periodicity of the unit cell is captured by Eq. (7). The velocity and concentration are equal along a pair of periodic boundaries (the subscripts 1 and 2 refer to a pair of periodic boundaries as indicated in Figure 5) while the gradients of velocity and concentration are equal in magnitude but opposite in sign because of the difference in unit outward normal directions for the two boundaries. The axial velocity field, v, is independent of axial distance, z, because of the well-developed velocity field and low mass transfer rate assumptions. Consequently, one can solve for v, substitute it into Eq. (2), and avoid solving the two equations simultaneously. Since an analytical solution cannot be found, a numerical approximation to the solution for v is obtained using the Finite Element method as described in the next section. If shell flow rate changes significantly due to flow across the fiber wall, one may still use the theoretical approach described here but must solve the conservation of momentum equation simultaneously with the conservation of mass equation - this greatly increases the computational complexity. Solving the conservation of mass equation, Eq. (2), for the concentration field allows one to calculate mass transfer coefficients. The effective, overall mass transfer coefficient, k, is given by: N
Z I (Cb -cw),Rao k(2n'RN ) i=1 2xRN
N
= ~, I2x D ( O C / O r ) ~ , i R d O i=1
(8)
where N is the number of fibers in the unit cell, Cb the mixing cup average bulk concentration given by: CVdA Cb = A
(9)
IvdA
A
and A the area available for flow. Eq. (8) can be written in terms of a local Sherwood number as: 13
Mass TransferIn Axial Flows ThroughRandomlyPacked Fiber
Bundles - L i p s c o m b
N
2Rk
Sh = ~ =
~,f2o~(a¢/a¢)w,iao
2 i=l
D
iv y, f2~r(c b _ c w l i d O
(10)
i=1
where ~ is the dimensionless radial coordinate (r/R). For a constant wall concentration boundary condition, Eq. (10) reduces to:
Sh e = 2 * 0¢/0.....~ Co
(11)
while for a constant wall mass flux boundary condition it reduces to: 2 Shf = (cb _cw )
(12)
where the overbar indicates the average value along the surface of each fiber in the unit cell. Values for the Sherwood number will be computed as a function of the Graetz number: (13)
Gz = Q/DL = Re*Sc*(xR/2L)*(1/~- 1)
where Q is the flow rate per fiber. In the large Graetz number limit (either high flow rates or short mass transfer regions), concentration boundary layers between adjacent fibers are much smaller than the fiber spacing. This limit is referred to as the entry mass transfer limit. In the small Graetz number limit (either low flow rates or long mass transfer regions), concentration boundary layers between adjacent fibers completely overlap. This limit is referred to as the well-developed mass transfer limit. 4.
NUMERICAL APPROXIMATIONS The Finite Element method [17] is used to obtain a numerical approximation to the governing conservation equations. To discretize the shell domain, V oronoi tessellation is used to surround each fiber in a unit cell with a polygon. The external boundary of these regions forms the boundary of the unit cell as illustrated in Figure 6. The region within each polygon is further discretized to form the finite element mesh. Note that the unit cell boundary consists of a series of line segments, instead of the four sides of the square unit cell in Figure 5, thus the periodic boundary consists of a sequence of line 14
Mass Transfer In Axial Flows Through Randomly Packed Fiber Bundles -
Lipscomb
segment pairs (see Figure 6). The use of this boundary facilitates the calculations, because one does not have to identify where fibers cross the boundary. The velocity field is calculated first from the Finite Element discretization of Eq. (1) and substituted into Eq. (2). Although the velocity field is independent of axial position, z, the concentration field is not and the Finite Element discretization of Eq. (2) leads to the following set of order differential equations: M -dc - + Kc = F dz
(14)
where c is now a vector containing the concentration at each node in the Finite Element mesh, M and K are coefficient matrices and F is a coefficient vector. lb
1~ \''~ S.
Figure 6. Use of Voronoi tessellation to generate the boundary of the unit cell and discretize the shell domain. Two pairs of periodic boundaries are indicated: (la, 2a) and (lb, 2b). A small portion of the actual mesh is also shown. These equations are solved by approximating dc/dz with a second-order, backward Finite Difference and marching along the z-direction to calculate the concentration field from the start to the end of the mass transfer region. A variable axial step size (Az) is used to speed up the calculations - the smallest values are used at the start of the mass transfer region where large concentration gradients exist. The calculations are stopped when the difference between the last two mass transfer coefficients is less than 0.001 in magnitude (approximately 0.1% change) or the bulk concentration is less than 10-6. One can avoid this computationally intensive solution for the concentration field as a function of z in the entry mass transfer limit. For large Graetz numbers, the relationship between Sh and Gz is given by [18]:
15
Mass Transfer In Axial Flows Through Randomly Packed Fiber Bundles -
Lipscomb
(15)
Sh = a G z 1/3
where the constant a depends on ~b and the wall boundary condition. As demonstrated previously [19], a can be calculated from the velocity field alone for both constant wall concentration and mass flux boundary conditions, so one does not have to obtain a numerical approximation to Eq. (2). In the welldeveloped mass transfer limit, the relationship between Sh and Gz is of the form
[2o]: (16)
Sh = fl
where the constant fl also depends on ¢~and the wall boundary condition. As demonstrated previously [21], flf (the subscript f indicates a constant wall mass flux boundary condition) can be calculated from that portion of the welldeveloped concentration field that is independent of z. One must obtain a numerical approximation to the well-developed portion of c, but marching along the z-axis is not required. Unfortunately, a similar solution does not exist for tic (the subscript c indicates a constant wall concentration boundary condition) and one must solve for the concentration field as a function of x, y, and z to evaluate Values for the Sherwood number calculated using each of these approaches are summarized in the next section. Note that one may use the marching solution procedure to obtain values for a and fl for both boundary conditions. The results obtained by marching are in excellent agreement with the results for ac, o?, and flfobtamed using the simplifications just described. Mesh refinement analyses are performed to evaluate discretization error. The number of nodes is increased by 50% until the calculated mass transfer coefficients change by less than 5% over the range of Graetz numbers considered. Additionally, mass transfer coefficients are calculated for regular arrays to validate the computational approach. 5.
RESULTS AND DISCUSSION Figure 7 illustrates the calculated variation of the local Sherwood number with Graetz number for triangular packings and a constant wall concentration boundary condition using the marching solution approach. The results are compared to those reported by Miyatake and Iwashita [22] for the equivalent heat transfer problem. The agreement between the two is very good and is comparable to that observed for square packings. We believe the minor differences that exist are due primarily to differences between Miyatake and Iwashita's actual results and the correlation they propose to represent the results; the correlation was used to generate Figure 7. 46
Mass Transfer In Axial Flows Through Randomly Packed Fiber Bundles Lipscomb -
Table 2 contains the values of a and fl calculated for triangular and square packings as a function of packing fraction and wall boundary condition. The computational approach used to evaluate each is indicated. Miyatake and Iwashita's results [22, 23] are also given; note that these values correspond to specific results they report and are not generated from the correlation they propose. The agreement between the two is excellent which helps validate the work summarized here. .
.
.
.
.
.
.
.
i
.
.
.
.
.
.
.
.
i
.
.
.
.
.
.
.
.
,
.
.
.
.
.
.
.
.
,
.
.
.
.
.
.
.
.
0.75 102
O.e3 0.40 0.23
Sh 10
.......
o
101
10 2
10 3
10 4
10 s
Gz
Figure 7. Variation of local Sherwood number with Graetz number for triangular fiber packings and constant wall concentration boundary condition. The solid lines are the results obtained using the marching algorithm described here and the dashed lines are the correlation proposed by Miyatake and Iwashita [22]. The values of ~ are indicated on the curves.
In-between the entry and well-developed mass transfer limits, the following expression provides a simple correlation of the relationship between Sh and Gz [24]:
Sh = (fl3 + ot3Gz~/3
(17)
Figure 8 compares the Sherwood numbers generated by this correlation and the actual calculated values for regular triangular arrays and a constant wall concentration boundary condition. Differences between the two are small except for the highest fiber packing, ~b= 0.75, which is outside the range of commercial interest; note that fiber packing ranges from 0.4 to 0.6 in most commercial modules. Consequently, mass transfer coefficients calculated for random packings will be reported in terms of a and fl and Eq. (17) will be used to calculate Sh. Figure 9 illustrates the differences between velocity fields calculated for a square packing and typical random packing with 0~=0.50. Regions of high flow where fiber packing is lowest are dearly visible in the randomly packed bundle. 17
Mass Transfer In Axial Flows Through Randomly Packed Fiber Bundles - Lipscomb
This flow channeling is undesirable and reduces mass transfer coefficients. Where fibers are dose, only a few contours are visible indicating low velocities and flows. Such regions are nearly stagnant and also undesirable. Table 2. Values of a and fl calculated for triangular and square packings. Values for a~ and af were taken from [ 19]. Values for/~ were taken from [21]. Values for tic were obtained using the marching algorithm described in the text. Triangular army - Constant wall concentration boundary condition ot~ This Work Miyatake-Iwashita [22] This Work 0.823 3.15 3.15 9.49 0.750 2.61 2.56 9.94 0.630 1.93 1.88 9.89 0.403 1.12 1.11 6.92 0.227 0.70 0.71 4.21 0.057 0.31 0.32 1.98
Miyatake-Iwashita [22] 9.52 9.97 9.89 6.86 4.17 1.96
Triangular army - Constant wall mass flux boundary condition a? This Work Miyatake-Iwashita [23] This Work 0.823 3.74 3.64 4.89 0.750 3.14 3.02 8.78 0.630 2.33 2.28 11.8 0.403 1.35 1.39 7.59 0.227 0.84 0.87 4.48 0.057 0.37 0.37 2.05
Miyatake-Iwashita [23] 4.92 8.75 11.6 7.49 4.41 2.04
Square array - Constant wall concentration boundary condition o~ This Work Miyatake-Iwashita [22] This Work 0.712 1.80 1.81 4.05 0.649 1.67 1.68 4.13 0.545 1.41 1.41 4.27 0.349 0.95 0.94 4.19 0.196 0.62 0.62 3.34 0.049 0.28 0.28 1.81
Miyatake-Iwashita [22] 4.07 4.17 4.33 4.21 3.29 1.82
Square array - Constant wall mass flux boundary condition o? ¢~ This Work Miyatake-Iwashita [23] This Work 0.712 2.07 2.06 2.29 0.649 1.97 1.92 3.12 0.545 1.70 1.63 4.42 0.349 1.14 1.14 4.98 0.196 0.75 0.77 3.68 0.049 0.34 0.35 1.89
Miyatake-Iwashita [23] 2.31 3.11 4.32 4.74 3.45 1.67
18
Mass Transfer In Axial Flows
Through
Randomly
Packed
Fiber Bundles
Lipscomb
-
Figure 10 illustrates the concentration field in the well-developed limit for the same randomly packed unit cell as shown in Figure 9. The concentration is nearly zero everywhere except in the lowest fiber packing region which corresponds to the highest flow region in Figure 9. Fluid residence time is lower in the low packing regions than in the high packing regions resulting in smaller concentration changes - one may envision a low packing region as a channel that allows fluid to "by-pass" the mass transfer region. The low packing regions dominate overall mass transfer performance because flow rates are much higher through them than high packing regions. ........
,
........
|
........
|
........
|
........ 0.75
102 0.63 0.40
0.23
Sh 101
.
.
.
.
.
.
.
.
.
.
.
.
.
.
.
.
.
.
........
i
104
.......
105
Gz
Figure 8. Comparison of the proposed local Sherwood number correlation (dashed lines) with calculated results (solid lines) for triangular packings and constant wall concentration boundary condition. The values of # are indicated on the curves.
F4 (a) (b) Figure 9. Contours of constant velocity in (a) regular square and (b) random fiber packings.
19
Mass Transfer In Axial Hows Through Randomly Packed Fiber Bundles - Lipscomb
Figure 10: Contours of constant concentration in a randomly packed bundle in the welldeveloped mass transfer limit for a constant wall concentration boundary condition. Table 3. Values of a and fl calculated for random packings. Random array - Constant wall concentration boundary condition a~ [19] Fiber Min. Max. Ave. Fiber Number Number 0.30 9 0.580 0.709 0.643 32 0.40 9 0.799 0.889 0.860 32 0.50 9 0.971 1.19 1.08 9 0.50 16 1.01 1.21 1.08 16 0.50 32 32 0.50 64 64 0.50 100 1.12 1.15 1.14 100 0.55 9 1.09 1.33 1.22 32 0.575 9 1.17 1.46 1.32 32 Random array - Constant wall mass flux af [19] Fiber Min. Max. Number 0.30 9 0.604 0.808 0.40 9 0.50 9 1.08 1.43 0.50 16 1.09 1.31 0.50 32 0.50 64 0.50 100 1.23 1.28 0.55 9 1.28 1.40 0.575 9 1.11 1.53
Min.
Max.
Ave.
0.484 0.480 1.11 0.914 0.694 0.758 0.676 0.713 0.615
0.840
0.623 0.918 1.71
1.23 2.83 2.08 1.27
1.27
1.19
0.968 0.987 0.800 0.928 0.904
]~ [21] Max. Min.
Ave.
0.185 0.329 0.225 0.153 0.165 0.142 0.175 0.154 0.098
0.213 0.355 0.460 0.342 0.331 0.247 0.252 0.197 0.144
1.20 0.908
1.20
boundary condition Ave.
Fiber Number 64 64 9 16 32 64 100 64 64
0.742 1.25 1.20 1.27 1.34 1.35
20
0.275 0.387 0.657 0.504 0.452 0.373 0.415 0.256 0.199
Mass Transfer In Axial Flows Through Randomly Packed Fiber Bundles - Lipscomb
Values calculated for a and fl for randomly packed bundles are summarized in Table 3. The maximum, minimum, and average values for at least five different unit cells are given. Table 3 also indicates the effect of unit cell fiber number. For 50% fiber packing, a and fl were calculated for unit cells containing 9, 16, 32, 64, or 100 fibers. An analysis of the variance (ANOVA [25]) indicates there is no significant difference between the results at the 95% confidence level if a sufficiently large number of fibers is used: at least 9 for c~c and o~, 32 for tic, and 64 for fl~ The results of this analysis were used to specify the fiber number for calculations with other ~ values. A comparison of Tables 2 and 3 indicates the detrimental effect of random fiber packing on mass transfer performance. Relative to a square packing, the random packing values for ac and af are-~20% lower. The decrease in the welldeveloped limit is much more dramatic: tic is 80% lower while flf is 90% lower. This detrimental effect is graphically illustrated in Figure 11 for ~=0.40 where the rapid decline in performance with decreasing Gz is readily apparent. 102
....................................
• .......
101
J J
Sh
J J J 10 0 ___.
1 0110_ 1
J
__-.-
10 °
101
10 2
10 3
04
GZ
Figure 11. Variation of local Sherwood number with Graetz number for regular square (solid line) and random (dashed line) fiber packings with ~=0.40 and constant wall concentration boundary condition. The following correlations accurately capture the dependence of a and fl on fiber packing fraction for randomly packed bundles - for a constant wall concentration boundary condition:
(18a) ( 8b)
ac = 2.44~- 0.097 tic = - 1 0 . 2 g + 9.92~- 1.43 and for a constant wall mass flux bounda~ condition:
21
Mass Transfer In Axial
Flows T h r o u g h
R a n d o m l y Packed Fiber Bundles -
Lipscomb
(19a)
off= 2.30~ + 0.061 f l f = - 8 . 1 5 g + 6 . 8 2 ~ - 1.09
(19b)
Substituting Eq. (18) and (19) into Eq. (17) gives a correlation that can be used to predict local shell-side mass transfer coefficients for randomly packed bundles as a function of G z and ~k (0.30< ~<0.575). The m e a n mass transfer coefficient, or equivalently Sherwood number is given by [26]:
Sh,,, = Lo
(20)
az
where the subscript m indicates mean value. Evaluating the m e a n Sherwood number in the entry and well-developed mass transfer limits leads to the following modification of Eq. (17) as a correlation for Shr,,:
+ 0 s ) oz] where a and fl are given by Eq. (18) and (19), respectively. The predictions of Eq. (21) for a constant wall concentration are compared to various literature correlations in Figure 12. Only those correlations 10 2
.
.
.
.
.
.
I
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.
.
.
.
.
101
Sh.
10 0
1 0 "1
101
.
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.
i
10 2
.
.
.
.
.
.
.
.
10 3
Gz
Figure 12. Comparison of calculated Sherwood numbers for randomly packed bundles with experimental correlations for 4~=0.50, Sc=1000, and 2R/L=0.003. The sources of the correlations are: * Costello et al. [10], 0 Gawronski and Wrzesinska [12], and small dot Wu and Chen [13]. The n represent theoretical predictions for a square packing and the line without symbols represents the random packing results both for constant wall concentration. valid for the chosen operating conditions are shown, and the predictions for a constant wall mass flux are just slightly higher than the constant wall
22
Mass Transfer In Axial Flows Through Randomly Packed Fiber Bundles - Lipscomb
concentration values. Although the experimental boundary condition is neither constant wall concentration or mass flux, one expects the theoretical predictions to provide bounds for the experimental values as observed for the lumen mass transfer coefficient. Clearly, the available correlations predict mean Sherwood numbers that are both much higher and much lower than theoretical predictions for either regular or random fiber packings. Why is this the case? One may attribute the positive (higher values) deviations to the presence of significant cross-flow effects that arise from either transverse flows engendered by non-parallel fibers or fluid distribution from the shell port (see Figure 2). As in heat exchange, concentration boundary layers do not increase as rapidly in cross-flow as in axial flow. The reduction in concentration boundary layer thickness increases mass transfer coefficients. One may attribute the negative (lower values) deviations to a combination of the other complexities of shell-side flow illustrated in Figure 2. First, poor fluid distribution from the shell inlet port may lead to higher flows in some regions and lower flows in others than one would expect if the axial pressure gradient was uniform throughout the bundle. Radial variations in the axial pressure gradient would be expected if the shell port and associated fluid distribution collar are not designed well: one would expect slightly higher axial pressure gradients closer to the port due to pressure drop through the distribution collar around the bundle. Evidence for higher flows in that portion of the bundle closest to the inlet port (see Figure 2) is given in the literature [7]. Poor packing of the fibers near the case would exacerbate this effect, since gaps between the bundle and case serve as conduits for fluid bypassing or channeling. Finally, the presence of large cross-flow regions may reduce the effective mass transfer coefficient due to poor utilization of the chemical potential driving force for transport. The lower driving force would be partially countered by the reduction in concentration boundary layer thickness discussed previously. Thus, mass transfer coefficients might decrease or increase depending on the extent and location of cross-flow regions. 6.
CONCLUSIONS Calculations of mass transfer coefficients for axial flows through randomly packed fiber bundles are presented. The results indicate randomness in fiber packing reduces mass transfer coefficients relative to structured, regular fiber packings. The reduction is most pronounced in the well-developed mass transfer limit, Gz
Mass Transfer In Axial Flows Through Randomly Packed Fiber Bundles - Lipscomb
regions due to shell port design or non-parallel fibers may lead to higher than expected mass transfer coefficients. Poor fluid distribution from the shell inlet port into the bundle or excessive cross-flow may lead to lower than expected values. Additional work is required, though, to quantitatively evaluate these effects. 7.
ACKNOWLEDGMENT The authors would like to acknowledge partial support of this work by the National Science Foundation through Grant CTS-9408414.
So A=
NOTATION 2 flow area, m 6,= dimensionless concentration C = concentration, mol/m 3 D = diffusion coefficient, m2/s vector from source term, generated by Finite Element discretization F= Gz = Graetz number (Q/DL) i= number of unit cell translations j = number of unit cell translations Jw = mass flux, mole/m:/s k= mass transfer coefficient, m/s K = coefficient matrix 1= unit vector L = module length, m M = coefficient matrix n= unit outward normal direction N = number of fibers in the unit cell P = pressure, Pa Q = flow rate, m3/s r= radial coordinate, m R = fiber radius, m Re = Reynolds number (2RVbp/,u) Sc = Schmidt number (/1/pD) Sh = Sherwood number (2Rk/D) dimensionless axial velocity V-axial velocity, m/s V = dimensionless coordinate x:" coordinate value, m X= y= dimensionless coordinate y= coordinate value, m z'-dimensionless axial coordinate Z = coordinate value, m 24
Mass Transfer In Axial Flows Through Randomly Packed Fiber Bundles - Lipscomb
At = p= 0=
8°1.
0= b= ¢=
f= m= w = X =
y=
coefficient in mass transfer coefficient correlation coefficient m mass transfer coefficient correlation fiber packing fraction fluid viscosity, Pa- s density, kg/m 3 angular coordinate dimensionless radial coordinate
Subscripts and superscripts initial value bulk value evaluated for a constant wall concentration evaluate for a constant wall mass flux mean value wall value x-coordinate value y-coordinate value
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[21 [31 [4]
[5] [6]
[7]
[8] [91 [10]
[111
J.T. Daugirdas, T.S. Ing, P.G. Blake, Handbook of Dialysis, 3rd Ed., Little Brown & Co., Boston, 2000. M. Cheryan, Ultmfiltration and Microfiltration Handbook, Technomic Publishing Co. Inc., Lancaster-Basel, 1998. A. Gabelman, S.T. Hwang, Hollow fiber membrane contactors, J. Membr. Sci. 159 (1999) 61. J. Henis, Commercial and practical aspects of gas separation membranes, pp. 441-512, Polymeric Gas Separation Membranes, D.R. Paul and Y.P. Yampol'skii, Eds., CRC Press, Boca Raton, 1994. L. Dahuron, E.L. Cussler, Protein extraction with hollow fibers, AIChE J. 34 (1988) 130. E.L. Cussler, Diffusion Mass Transfer in Fluid Systems, Cambridge University Press, Cambridge, 1997. A. Frank, G.G. Lipscomb, M Dennis, Visualization of concentration fields in hemodialyzers by computed tomography, J. Membr. Sci. 175 (2000) 239. M.C. Yang, E.L. Cussler, Designing hollow fiber contactors, AIChE J. 32 (1986) 1910. R. Prasad, K.K. Sirkar, Dispersion-free solvent extraction with microporous hollowfiber modules, AIChE J. 34 (1988) 177. M.J. Costello, A.G. Fane, P.A. Hogan, R.W. Schofield, The effect of shell side hydrodynamics on the performance of axial flow hollow fiber modules, J. Membr. Sci. 80 (1993) 1. R.M.C. Viegas, M. Rodriguez, S. Luque, J.R. Alvarez, I.M. Coelhoso, J.P.S.G. Crespo, Mass transfer analysis of Wilson-plot methodology, J. Membr. Sci. 145 (1998) 129.
25
Mass Transfer In Axial Flows Through Randomly Packed Fiber Bundles -
[12] [13] [14]
[15] [16]
[17] [18] [19] [20] [21]
[22]
[:3] [24]
[25] [26]
Lipscomb
R. Gawronski, B. Wrzesinska, Kinetics of solvent extraction in hollow-fiber contactors, J. Membr. Sci. 168 (2000) 213. J. Wu, V. Chen, Shell-side mass transfer performance of randomly packed hollow fiber modules, J. Membr. Sci. 172 (2000) 59. F. Lipnizki, R.W. Field, Mass transfer performance for hollow fibre modules with shell-side feed flow: using an engineering approach to develop a framework, J. Membr. Sci. 193 (2001) 195. A.S. Sangani, Private communication. J.D. Rogers, R.L. Long, Modeling hollow fiber membrane contactors using film theory, Voronoi tessellations, and facilitation factors for systems with interface reactions, J. Membr. Sci. 134 (1997) 1. J.N. Reddy, D.K. Gartling, The Finite Element Method in Heat Transfer and Fluid Dynamics, CRC Press, Boca Raton, FL, 1994. L.G. Leal, Laminar Flow and Convective Transport Processes, ButterworthHeinemann, Boston, 1992. L. Bao, B. Liu, G.G. Lipscomb, Entry mass transfer in axial flows through randomly packed fiber bundles, AIChE J. 45 (1999) 2346. W.M. Deen, Analysis of Transport Phenomena, Oxford University Press, New York, 1998. L. Bao, G.G. Lipscomb, Well-developed Mass Transfer in Axial Flows Through Randomly Packed Fiber Bundles with Constant Wall Flux, Chem. Eng. Sci. 57 (2002) 125. O. Miyatake, H. Iwashita, Laminar flow heat transfer to a fluid flowing axially between cylinders with a uniform surface temperature, Int. J. Heat Mass Transfer 33 (1990)417. O. Miyatake, H. Iwashita, Laminar flow heat transfer to a fluid flowing axially between cylinders with a uniform wall heat flux, Int. J. Heat Mass Transfer 34 (1991) 322. H. Kreulen, C.A. Smolders, G.F. Versteeg, W.P.M. van Swaaij, Microporous hollow fiber membrane modules as gas-liquid contactors. Part 1. Physical mass transfer processes, J. Membr. Sci. 78 (1993) 197. E.A. McBean, F.A. Rovers, Statistical Procedures for Analysis of Environmental Monitoring Data & Risk Assessment, Prentice Hall, Upper Saddle River, NJ, 1998. W.M. Kays, M.E. Crawford, Convective Heat and Mass Transfer, McGraw-Hill, New York, 1980.
26
New Insights into Membrane Science and Technology: Polymeric and Biofunctional Membranes D. Bhattacharyya and D.A. Butterfield (Editors) © 2003 Elsevier Science B.V. All rights reserved.
Chapter 2
Fouling phenomena during microfiltration: effects of pore blockage, cake filtration, and membrane morphology Andrew L. Zydney *+, Chia-Chi Ho s, Wei Yuan # +Department of Chemical Engineering, The Pennsylvania State University, University Park, PA 16802 SDepartment of Chemical Engineering, University of Cincinnati, Cincinnati, OH 45221 #Innovasep Technologies, 420 Maple Street, Marlborough, MA 01752 * Corresponding Author, Phone: 814-863-7113, Fax: 814-865-7846, e-mail:
[email protected] 1.
INTRODUCTION
Fouling remains a major problem in many membrane filtration processes. Particulate matter can deposit on or within the membrane pore structure, significantly increasing the overall hydraulic resistance to flow. This can lead to dramatic reductions in the filtrate flux, ultimately requiring mechanical/chemical cleaning or complete replacement of the filter media. The reduction in filtration flux (or velocity) during constant pressure filtration is typically described using one of the classical fouling models. These blocking laws were originally developed by [1] by assuming that the filtering medium was composed of a parallel array of capillary cylindrical pores of constant and equal diameter. The complete blocking law assumes that particulate matter deposits on the membrane surface and completely blocks (or seals) the capillaries. The standard blocking law assumes that the particles deposit throughout the pore volume, causing a uniform constriction m the pore radius. It is also possible for the particles to form a growing deposit or cake layer on the external surface of the filter. The cake filtration model is developed by assuming that the rate of cake growth is uniform across the membrane surface and is directly proportional to the filtration velocity. Although these fouling models have been used quite extensively to analyze filtration data, the assumptions that underlie these models are almost never satisfied in actual practice. Most filtration media have a highly irregular pore structure formed by the void space between the polymer fibers or globules that make up the membrane or filter. The effective size of these pores can vary dramatically both laterally and with depth through the membrane filter. In
27
Fouling Phenomena During Microfiltrafion: Effects Of Pore Blockage, Cake Filtration, And Membrane Morphology - Zydney
addition, particle deposition on the upper surface of such a membrane will be unable to completely block or seal the pores since the fluid can flow under and around any surface blockage due to the highly interconnected nature of the pore structure. A cake layer may form on the external surface of these membranes, but the cake growth rate is likely to be non-uniform, with those regions of the filter that were fouled first having the thickest cake layer. Finally, the actual fouling in any given system is likely to be due to a combination of mechanisms, with pore blockage and cake growth occurring simultaneously during the filtration process. Researchers [1-3] showed that the cake filtration and blocking laws could all be written in a common mathematical form as: d:t- k(dt n dV2 dV )
(1)
where t is the filtration time and V is the total filtered volume. The exponent n characterizes the filtration model, with n = 0 for cake filtration, n = 3/2 for standard blocking, and n = 2 for complete blocking. Hermans and Bred6e [1] also coined the term "intermediate blocking" to describe the filtration behavior when n = 1, although no physical interpretation of this situation was provided. Hermia [4] subsequently showed that this intermediate blocking law could be derived from the complete blocking model by allowing for the superposition of deposited particles on the filter surface. More recently, derived the governing filtration equations for 11 different values of n ranging between 0 and 2, including n = 1/4, 1/3, 1/2, 2/3, 3/4, 5/4, and 4/3 were derived [5], although no physical interpretation was provided for these model equations. Equation (1) has often been used to analyze experimental results during membrane filtration; however, much of the flux decline data obtained in these studies is difficult if not impossible to explain using this simple theoretical framework. For example, analyzed filtrate flux data during microfiltration of bovine serum albumin (BSA) through track-etch polycarbonate membranes with straight-through cylindrical pores were analyzed [6], and these researchers found a distinct maximum in the plot of dZt/dV2 versus dt/dV. The slope on the log-log plot thus becomes negative at large dt/dV, corresponding to a negative value of n in Eq. (1). Similar behavior was seen for BSA filtration through nitrocellulose microfiltration membranes having a more tortuous and irregular pore structure [7]. In this case, the value of the slope on the log-log plot decreased throughout the filtration, going from a value of more than 6 at the start of the experiment to a value less t h a n - 2 at long filtration times. Such extreme values of n have never been explained, nor has there been any quantitative explanation for the large variation in n observed over the course of the filtration. This type of unusual filtration behavior is in no way limited to solutions of proteins like bovine serum albumin. For example, studied the flux decline during filtration of 28
Fouling Phenomena During Microfiltrafion: Effects Of Pore Blockage, Cake Fillrafion, And Membrane Morphology - Zydney
a sodium acetate suspension in methanol through a kaolin membrane preformed on a stainless steel support was reported [8]. Data were analyzed using Eq. (1), with the calculated values of the slope being consistently less than zero. Values as small as -15 were reported for experiments performed at high temperature and pressure. This type of filtration behavior was denoted as "decelerating resistance" [8], although no clear physical explanation was provided for this phenomenon. We have recently developed a new approach for describing the flux decline during membrane microfiltration that addresses many of the shortcomings in the classical fouling models and is able to explain much of the unusual flux decline phenomena reported in the literature. The following section of this Chapter describes the basic principles for this combined pore blockage - cake filtration model, including the effects of the underlying pore morphology on the flux decline. Section 3 examines the behavior of this theoretical model, using experimental data from the literature to help illustrate many of the key phenomena. The final section summarizes the key results and discusses the implications of this new understanding of membrane fouling for the design and operation of improved filtration processes. 2.
PORE B L O C K A G E - CAKE FILTRATION MODEL
Although fouling can occur by specific chemical interactions between feed components and the membrane material, our focus is on the fouling associated with the deposition of particulate matter on the upper surface of the membrane filter. This physical situation will certainly be appropriate for the filtration of suspensions in which the particles are larger than the membrane pore size. In addition, several studies have demonstrated that fouling during protein microfiltration occurs primarily by the physical deposition of large protein aggregates on the external membrane surface [9]. This type of aggregate deposition is also the primary cause of humic acid fouling during microfiltration [10]. The rate of surface coverage (pore blockage) is assumed to be proportional to the mass flow rate of large particles/aggregates to the open (unfouled) surface of the membrane: dA
open = _ aQopenCb dt
(2)
where Qopenis the volumetric flow rate through the open pores and Cb is the bulk concentration of large aggregates, o~ is a pore blockage parameter and is equal to the membrane area blocked per unit mass of aggregates convected to the membrane. Cb depends upon the properties of the protein and buffer
29
Fouling Phenomena During Microfillration: Effects Of Pore Blockage, Cake Filtration, And Membrane Morphology - Zydney
solution. Unlike most prior work, we allow for the possibility that the blocked pores may be partially permeable to filtration, e.g., the fluid may be able to flow through the interstitial space between individual macromolecules in the aggregate. As the membrane surface becomes more heavily fouled, the large particles or aggregates will also begin to deposit directly on the fouling layer, increasing its hydraulic resistance to flow. This is exactly what occurs in the classical cake filtration model, although in this case the cake growth is assumed to occur simultaneously with the coverage of the remaining open area of the membrane. The rate of deposit growth over each blocked region of the membrane is assumed to be proportional to the particle mass flux to that particular region: dm dt
(3)
P = Obloeked - J b ~ k ) f ' b
where mp is the protein mass deposited per unit membrane area and Jb,ck accounts for any back particle flux associated with shear induced diffusion, inertial particle lift, surface migration, or long-range repulsive interactions between the particles and cake layer [ 11 ]. C'b is the concentration of particulate matter that adds to the growing deposit. C'b can be greater than Cb if the deposit is able to retain a greater fraction of the feed stream during the filtration pr'ocess. In order to integrate Eqs. (2) and (3), we first need to develop expressions for the volumetric flow rate or filtrate flux (J = Q/A) through the open and blocked pores. These flow rates will not only be a function of the extent of membrane fouling, they will also depend upon the underlying pore structure and morphology of the membrane. In the following sub-sections we consider three distinct membrane morphologies: an isotropic membrane with noninterconnected (straight-through) pores, a composite membrane in which the upper ("skin") layer has non-interconnected pores and the lower layer (substructure) has very highly interconnected pores, and a homogeneous membrane with an interconnected pore structure throughout the membrane. 2.1
Non-Interconnected Pore Structure For a membrane with non-interconnected pores, the volumetric filtrate flow rates through the open and blocked pores are given as:
AP Qown - t.tRm Aopen
(4)
30
FoulingPhenomenaDuringMierofiltration:EffectsOfPoreBlockage,CakeFiltration,And Membrane Morphology - Zydney
A
Qblo~k~d
=
blocked
J'O ' ( R m + R p o + R ' m p
)dAbloeked
(5)
where ~t is the solution viscosity, Rm is the membrane resistance, Rpo is the initial resistance of the first particle/aggregate, and R' is the specific resistance of the particle deposit or cake layer on the membrane surface. Eq. (5) is expressed as an integral over the blocked area to account for the spatial inhomogeneity in the thickness (or mass) of the protein deposit arising from the time-dependent blockage of the membrane surface. Thus, the protein deposit on a given region of the membrane only grows over the time interval t-tp where tp is the time at which that particular region was first covered or blocked by a protein aggregate [12,13]. Those regions of the membrane surface that have been blocked most recently will have the smallest values of mp and thus the greatest local filtrate flux. Equations (2) - (5) can be integrated numerically over time using the approach presented by Ho and Zydney [12,13]. It is also possible to integrate these equations analytically by assuming that the mass of the protein deposit is approximately uniform over the membrane surface at its maximum value [12-14]. The net result is: Q - ex Qo
-
+ ~Rm
1 - ex Rm + Rpo + R' mp
-
t ~R m
(6)
with:
J back
~baek
/~
-- ~ b a e k
(Rm + Rp-~)
]
(7)
The first term in Eq. (6) is equivalent to the classical pore blockage model and gives a simple exponential decay in the volumetric flow rate. At long times, the volumetric flow rate is dominated by the second term and is thus proportional to the ratio of the membrane resistance to the total resistance. The mass of the cake layer is a complex fimction of time, although it is relatively easy to show that Eq. (7) reduces to the form given by the classical cake filtration model in the limit of Jback= 0 [ 14]. 2.2
Composite Membrane Structure The situation is more complex for a composite membrane due to the flow through the interconnected pores in the membrane substructure (Figure 1). The flux through the open and blocked pores in the upper layer of the membrane are expressed in terms of the pressure drop across the upper layer of the composite 31
Fouling Phenomena During MicrofUtralion: Effects Of Pore Blockage, Cake Filtration, And Membrane Morphology - Zydney
Figure 1" Schematic drawing of flow through a composite membrane structure. structure (AP1) using Eqs. (4) and (5) but with AP replaced by AP1. The total pressure drop across the composite membrane is expressed in terms of the total filtrate flow rate as: AP = AP1 +
gQRsub
(8)
A
where Rsub is the resistance of the membrane substructure. The area A in the denominator of Eq. (8) is the total membrane area since the effects of any pore blockage on the surface of the upper layer are "lost" in the bottom layer due the lateral fluid flow. This is discussed in more detail elsewhere [ 15]. Equations (4), (5), and (8) can be combined to evaluate the filtrate flux through the open and blocked pores as:
J open _ R m(Rm + R~ub) + (Rw + R' mp )(R m + R~ub) J o - Rm (Rm + R~ub) + (Rpo + R' m p)(R m + XR~ub) J blocked
Rm(Rm+ R~ub)
J0
Rm (Rm + Rsub)-b (Rpo + R' mpXRm -b XRsub )
(9)
(10)
where X = Aop~/A is the fraction of open pores at any time t. Note that Joppa actually becomes greater than J0 as the pores become blocked, i.e. as X decreases. This occurs because the fluid flow through the composite membrane is shunted away from the blocked pores and through the open pores. This effect is not seen in a homogeneous (single layer) membrane (i.e. when R~ub- 0), since there is no opporttmity for the fluid to flow laterally as it moves through such a pore structure. Equations (2) and (3) can be integrated numerically over time, with the filtrate flux through the open and blocked pores given by Eqs. (9) and (10) [15]. It is also possible to develop an analytical solution to these equations [15]. In this case the open pore area is evaluated iteratively as an implicit function of
32
Fouling Phenomena During Microfiltration: Effects Of Pore Blockage, Cake Filtration, And Membrane Morphology - Zydney
time, with the mass of the deposit expressed analytically in terms of the fractional open pore area.
2.3
lsotropic Membrane Structure Particle deposition on the upper surface of an isotropic membrane will only disturb the fluid flow over a small penetration depth into the pore structure since the fluid can flow laterally through the interconnected pores as it percolates through the membrane (Figure 2). The size of the surface blockage and the relative ease of lateral flow within the pore structure will determine the actual depth of the penetration.
....
Figure 2: Leftpanel: schematicof flow percolation through isotropic membrane Right panel: schematicof upper surface of partially fouled membrane In order to evaluate the filtrate flux through such a membrane structure, we assume that particle/aggregate deposition occurs uniformly and randomly over the upper surface membrane, giving the situation shown schematically in the fight hand panel of Figure 2. At low surface coverage the flow through this membrane can be described by a Krogh-cylinder model, with the surface blockage located at the center of a larger cylindrical region that sub-divides the membrane. The radius of this outer cylinder is directly related to the extent of surface blockage:
1--
Aopen A
m
/ 12 rbl 3eka~e
(11)
\ r~,,lindo~
As the membrane becomes more heavily fouled, the fluid will flow primarily through the gaps (or holes) between the fouled regions on the membrane surface. In this case, the membrane can be described using a similar model with the blockage occupying an outer annular region around a central void [ 16,17]. The velocity profiles within the porous membrane are described by Darcy's law:
33
FoulingPhenomenaDuringMicrofiltration: Effects OfPoreBlockage,CakeFiltration,AndMembraneMorphology- Zydney
c~ Vr = - K r
m
6].
c~ Vz = - K z - -&
(12)
where Kr and 1(7. are the Darcy permeabilities in the radial and transverse directions, respectively. A membrane with non-interconnected (straightthrough) pores would have Kr = 0 while a truly isotropic membrane would have Kr = Kz. The velocities Vr and Vz must satisfy the continuity equation within the membrane, leading to the following second-order partial differential equation for the pressure:
~P
+
~2
7
-0
(13)
where the ratio of the permeabilities (Kd~Kz) provides a quantitative measure of the ease of lateral flow. A novel experimental approach to evaluate this permeability ratio, which is really just a measure of the extent of pore connectivity, based on the relative fluid flow rate through a membrane that is partially covered with an impermeable tape was developed [ 12,13]. Equation (12) can be solved numerically for the pressure profiles within the homogeneous membrane. Symmetry conditions are applied in the radial direction at the center and outer edge of the Krogh cylinder. The boundary condition at the downstream surface of the membrane (z = ~Sm)is simply P = Pfiltr.te. A split boundary condition is applied on the upper surface to account for the hydraulic resistance of the deposit covering the blocked region of the membrane:
z-O" z=0:
Pf~ed -- P~=0
Vz--Kz"& =g(Rpo +R'mp) forO___r_
for rbloekage < r ___l'cylinder
(15)
The filtrate flux through the open and blocked regions are then calculated directly from the pressure profiles by numerical integration of the Darcy expression for Vz over the membrane area:
reylinder
_ J open
-1
.
/~/
[.2~rKz
'¢g~cylinder ~ 2 " - ~blockage)rblockage'I
dr z=O
34
(16)
FoulingPhenomenaDuringMicrofiltration:Effects Of Pore Blockage,CakeFiltration,And MembraneMorphology- Zydney rbloekage
Jblockea --
dr
2 ]/rblockage 0
z=0
(17)
The volumetric flow rate is then evaluated as a function of time by numerical integration of Eqs. (2) and (3) with the flux through the open and blocked pores given by Eqs. (16) and (17). 3.
F O U L I N G D A T A A N D ANALYSIS
The effects of the membrane morphology on the flux decline during the constant pressure filtration (AP = 14 kPa = 2 psi ) of 2 g ~ solutions of the protein bovine serum albumin (BSA) are shown in Figure 3. These BSA solutions contained 0.03 % protein aggregates, thus Cb and C'b in Eqs. (2) and (3) were both 0.6 mg/L. The data were obtained with three different membrane morphologies. The polycarbonate (PCTE) membrane is made by bombardment and etching of a polycarbonate film, resulting in a membrane with straightthrough (non-interconnected) nearly cylindrical pores. The polyvinylidene fluoride (PVDF) membrane is made by immersion casting and has a highly irregular and interconnected pore structure defined by the void space between polymer fibers. The PCTE/PVDF is a composite membrane formed by placing a PCTE membrane directly on top of an isotropic PVDF membrane. The data are plotted as the normalized flux (J/J0) as a function of the cumulative filtrate volume, where J0 is the initial flux through the membrane (values summarized in Table 1). The PVDF membrane has a smaller initial flux than the PCTE membrane due to its much greater thickness. The composite membrane has the smallest initial flux due to the additive nature of the resistances of the two membranes in series. Jo (m/sX 10 4)
Thickness (~Lm)
¢z (me k g -1 x 10 4)
(m -1 x 10 -11)
R' (m k g -1 x 10 -is)
PCTE
4.0
10
1.4 4-0.1
4.0 4- 0.2
8.0 4- 0.7
PVDF
2.2
120
0.7 4-0.1
4.0 4- 0.2
8.0 4- 0.7
Membrane
Rpo
PCTE/PVDF 1.7 130 1.1 4- 0.1 4.0 4-0.2 8.7 4-0.7 Table 1: PhysicalProperties and Fouling Parameters for Different Membranes The rate of flux decline for the PCTE membrane is very rapid, with the flux declining to less than 10% of its initial value within the first 60 mL of filtration. In contrast, the flux through the isotropic PVDF membrane remains greater than 80% of its initial value over the first 100 mL of filtration. The net result is a significantly greater capacity for the PVDF membrane, even though
3S
Fouling Phenomena During Microfiltration: Effects Of Pore Blockage, Cake Filtration, And Membrane Morphology - Zydaey
that membrane has a much lower initial flux. For example, the cumulative filtrate volume obtained before the flux drops below J = 4 x 10-5 m/s is about 60 mL for the PCTE membrane compared to more than 180 mL for the PVDF. The addition of the PVDF membrane beneath the PCTE membrane also reduces the extent of flux decline compared to that for the PCTE membrane alone, although the fouling remains much more dramatic than that for the PVDF membrane. Note that these differences in flux decline were not due to differences in the total pore volume as SEM micrographs demonstrated that all of the fouling in these experiments occurred on the upper surface of the membranes. The much smaller flux decline seen with the isotropic PVDF membrane is a direct result of the highly interconnected pore structure of this membrane. Aggregate deposition on the surface of the PCTE membrane leads to complete blockage of the pores, although the blocked pores do allow some fluid leakage through the interstices within the protein aggregate. In contrast, aggregate deposition on the surface of the PVDF membrane only disturbs the fluid flow over a small penetration depth into the membrane, with the fluid simply flowing under and around the surface blockage (Figure 2). The flow in the composite membrane is somewhat more complex, with the surface blockage effectively sealing the pores in the upper (PCTE) layer while the entire (PVDF) substructure remains available for flow. In this case, the fouling reduces the pressure drop across the lower layer of the membrane, leading to an increase m AP1 and a corresponding increase in the flux through the open pores [15]. This increase in Jo~ partially offsets the reduction in the area of the open pores, thus reducing the extent of flux decline compared to that for the PCTE membrane alone. The solid curves m Figure 3 are model calculations developed by numerical solution of Eqs. (2) and (3). The filtrate fluxes through the open and blocked pores were given by the appropriate equations for a membrane with straight-through non-interconnected pores (Eqs. 4 and 5), for a composite membrane structure (Eqs. 9 and 10), and for an isotropic membrane with highly interconnected pores (Eqs. 16 and 17). In each case the best fit values of the model parameters were determined by minimizing the sum of the squared residuals between the filtrate flux data and the model calculations using the method of steepest descent, with the results summarized in Table 1. Jbackwas set equal to zero for these calculations, although a slightly better overall fit to the data could be achieved using a small positive value for this parameter. Additional details on the model calculations are provided elsewhere [12,13,15]. The fairly sharp change in slope for the PVDF membrane around V = 180 mL occurs when the membrane surface is nearly completely covered by protein aggregates and reflects the transition from a pore blockage to a cake filtration mechanism. The model calculations are in very good agreement with the experimental data for all 3 membranes, demonstrating the ability of this simple
36
Fouling Phenomena During Microfillration: Effects Of Pore Blockage, Cake Filtration, And Membrane Morphology - Zydncy
O PCTE
~
0.8
PVDF PCTE/PVDF
0.6
0.4
r-! I-I
0.2
0
50
100
150
290
250
Filtrate Volume. V (mL) Figure 3: Effects of membrane morphology on the flux decline during the constant pressure filtration of 2 g/L solutions of BSA. Data are shown for the polycarbonate (PCTE), polyvinylidene fluoride (PVDF), and composite PCTE/PVDF membranes.
theoretical framework to describe the flux decline behavior of porous membranes with very different underlying pore morphologies. The best fit values of Rpo and R' for the PCTE, PVDF, and PCTE/PVDF membranes were essentially identical, with the actual values in good agreement with independent estimates of these fouling parameters obtained from measurements of the hydraulic resistance and mass of the protein deposits [16,17]. The values of the pore blockage parameter, (z, show some differences between the three membranes which is likely due to small batch-to-batch variations in the properties of the BSA solutions. The dashed curve in Figure 3 shows the model calculations for the PCTE membrane evaluated using the simple analytical solution (Eqs. 6 and 7) using the same parameter values given in Table 1. This approximate solution is in excellent agreement with both the experimental data and the full numerical results, even though the approximate solution was developed by assuming that the mass of the protein deposit is uniform across the membrane surface at its maximum value. This effect is shown more explicitly in Figure 4, in which the
37
Fouling Phenomena During Microfillralion: Effects Of Pore Blockage, Cake Filtration, And Membrane Morphology - Zydncy
calculated values of the deposit mass per unit membrane area are plotted as a function of the fractional surface coverage of the membrane at four different filtration times. 30 t = 32 min
25 20
15
t = 16 min
10 t = 8 min t=4
0
0.2
min
0.4
0.6
0.8
1.0
Fractional Blocked Area, Ab~oCkedA Figure 4: Model calculations for the deposit mass over the fouled region of the membrane. Solid curves are full numerical solution. Dashed curves are approximate analytical solution.
The dashed curves show the calculations using the approximate analytical solution, in which case the deposit mass is uniform across the fouled region of the membrane at its maximum value. The solid curves are the calculations using the full numerical solution and show the variation in deposit mass over the fouled region of the membrane. Even at long filtration times, the deposit mass remains relatively uniform over the membrane surface due to the inherent "selfleveling" character of the fouling process. Those regions of the membrane that are fouled first have the smallest filtrate flux, leading to the slowest rate of cake growth. Those regions that were fouled most recently have the thinnest cake but the greatest rate of cake growth. The net result is that the average protein layer resistance over the fouled region of the membrane at t = 32 mm is only 6% smaller than the maximum resistance, with this difference decreasing to less than 3% after 100 mm of filtration. This self-leveling phenomenon is the reason why the simple approximate solution provides such accurate predictions for the filtrate flux as seen in Figure 3.
38
Fouling Phenomena During Microfiltration: Effccta Of Pore Blockag©, Cake Filtration, And M©mbran¢ Morphology - Zydacy
Additional insights into the fouling behavior can be obtained by analyzing the flux decline data using the functional form suggested by Eq. (1). The required derivatives can be calculated directly from the filtrate flux data versus cumulative filtrate volume data in Figure 1 as: dt
m
dV dZt dV 2 -
1
(18)
JA 1 (dJ) j2 A ~,~'~)
(19)
where dJ/dV was evaluated numerically [16,17]. The results are shown in Figure 5, with the solid curves representing the model calculations developed using the equations in Section 2 using the same values of the model parameters shown in Table 1. The model calculations are again in very good agreement with the data for all three membranes. The small discrepancy at large dt/dV for the PVDF membrane is due to the uncertainties associated with the transitions from pore blockage to cake filtration and from flow around a centrally located blockage to flow through a cylindrical region between aggregates. At low dt/dV (corresponding to high filtrate flux), the data for the PCTE membrane yield a linear relationship with slope approximately equal to 2.0, which is consistent with the pore blockage mechanism. In contrast, the data for the PVDF and the composite PCTE/PVDF membranes show much steeper slopes at the start of the filtration. All three membranes show a distract maximum in d2t/dV 2 at an intermediate dt/dV. Others [6,7] also have seen this maximum in d2t/dV 2, a result that has never been explained previously and is completely inconsistent with the behavior predicted by the classical fouling models. This behavior is very accurately described by the pore blockage - cake filtration model, with the transition between these two fouling mechanisms causing a sharp reduction in the rate of flux decline leading to the observed maximum in d2t/dV 1. The transition between the pore blockage and cake filtration fouling mechanisms can be seen more dearly by numerically evaluating the slope of the d2t/dV 2 versus dt/dV data in Figure 5. The results are shown in Figure 6 as a function of the filtration volume. The large scatter in some of the data is due to the errors involved in the numerical differentiation of the raw data. The solid curves again represent the model calculations using the parameter values given in Table 1. The slope for the PCTE membrane begins around 2, and remains nearly constant at that value over the first 60 mL of filtration. This behavior arises from dominance of the
39
Fouling Phenomena During Microfiltration: Effects Of Pore Blockage, Cake Filtration, And Membrane Morphology
- Zydncy
1013
l"1
1012 E
N~
1011 0 PCTE r"l PVDF A PQ,TE/PVDF 10lo
. . . .
I
108
10 7
109
dt/dV (s m "3) Figure 5: Flux decline analysis for BSA filtration through the PCTE, PVDF, and composite PCTE/PVDF membranes. Experimental data are from Figure 3. Solid curves are model calculations.
pore blockage mechanism; the slightly smaller value of the slope compared to that for the classical pore blockage model is due to the small leakage flow through the interstitial space within the protein aggregates. The transition from pore blockage to cake filtration occurs when the membrane surface is nearly completed covered with protein aggregates, and it is marked by a rapid reduction m slope to a value less than -6. The slope then increases and attains a value of zero for V > 85 mL, consistent with the dominance of the cake filtration mechanism at long times. The slope for the composite membrane beings around 6 and decreases monotonically during the start of the filtration process. This reduction in slope reflects the shunting of the fluid flow towards the open regions of the membrane. There is again a sharp decrease m slope around V = 100 mL due to the transition between a pore blockage dominated fouling to a cake filtration mechanism. The initial slope for the isotropie PVDF membrane is around 9 and decreases in value throughout the filtration run. Similar extreme values of the slope have also been reported previously [7,8], but they never have been explained fundamentally. Our new theoretical framework provides a good qualitative prediction of the observed behavior,
40
Fouling Phenomena Dunng ~.~crofiitration: Effects Of Pore Blockage, Cake Filtration, And Membrane Morphology -Zydney
10
,~~,
8 c:
6
a-
4
.=-
2
O
0
..O
.,--.,.
u,.,
DF
O
O PCTE r l Pvog ~ PCTE/PVDF
Z~ A
%~'~A E]
-2
PCTE/PVDF
,
0
O
E!
-4.
-8
~,n ,-,
PCTEt~:~ #~)O O
II--
O
,.~ '
I
50
,
,
100
I
150
,
I
200
n ,
250
Cumulative FiltrateVolume, V (mL) Figure 6: Slope determined from the plot of d2t/dV 2 versus dt/dV for the PCTE, PVDF, and composite PCTE/PVDF membranes. Experimental data are from Figure 3. Solid curves are model calculations. with the quantitative discrepancies seen in Figure 6 arising from the multiple derivatives that have to be taken to evaluate the slope. Thus, there is no need to invoke any new or complex fouling phenomena to explain these experimental observations. The key is the proper inclusion of both the pore blockage and cake filtration behavior and the proper analysis of the flow properties of membranes with different underlying pore morphologies. Although the experimental data presented in Figures 3 - 6 were all obtained with bovine serum albumin, the pore blockage- cake filtration model can also explain the flux decline behavior seen in a wide variety of filtration systems. For example, Figure 7 shows results for the constant pressure filtration of BSA, lysozyme (a 14 kD molecular weight protein), and a soil-based humic acid through the 0.2 gm pore size polycarbonate track-etched membranes Additional details on the protein and humic acid experiments are provided elsewhere [14,18], respectively. In each case, the derivatives were evaluated numerically directly from data for the filtrate flux as a function of the cumulative filtrate volume. The solid curves are calculations using the simple approximate solution (Eq. 6) with the best fit values of the parameters determined by comparison of the model and data. Jbackwas set equal to zero for both BSA and lysozyme, while the best fit value of Jbackwas 5.3 x 104 m s1 for
41
Fouling Phenomena During Mierofiltration: Effects Of'Pore Blockage, Cake Filtration, And Membrane Morphology - Zydney
lo 3
A ,,f,
.
.
.
.
.
.
.
.
i
.
.
.
.
.
.
.
.
i
.
.
.
.
.
.
.
.
,~
1012
E m
"o ol
"o 1011
LJ 10lo 10 6
113_ t.yso me
~
..~.. Hum.ic~.Ac.id
10 7
10 8
10 9
dt/dV (s m "3) Figure 7: Flux decline analysis for the constant pressure filtration of solutions of BSA (2 g/L), lysozyme (2 g/L), and a soil-based humic acid (2 ppm) through polycarbonate track-etch membranes. Solid curves are model calculations.
the 2 ppm humic acid. The initial slope for all three systems was approximately equal to two, which is again consistent with the dominance of the pore blockage phenomenon at the start of the filtration. The flux decline data for the humic acid solution show a distinct maximum in d2t/dV2 reflecting the transition between the pore blockage and cake filtration mechanisms. At long filtration times, dZt/dV2 for the humic acid solution approaches zero as the flux approaches its steady-state value J = Jb,ck. This sharp decline in d2t/dV2 provides a convenient diagnostic for determining the presence of a steady-state flux in any given filtration process. The lysozyme data show a much smoother transition from the pore blockage to cake filtration mechanisms, with no distinct maximum observed in the dZt/dV2 data. This occurs because of the greater rate of cake growth relative to pore blockage for the lysozyme [ 18]. 4.
SUMMARY
Membrane fouling remains a major problem in applications of microfiltration for particle removal, sterile filtration, and clarification. The theoretical framework presented in this chapter provides a completely new approach for the analysis and interpretation of flux decline data obtained in these membrane systems. The model explicitly accounts for fouling due to surface (pore) blockage and cake filtration, with the cake forming over those
42
Fouling Phenomena During Microfiltration: Effects Of Pore Blockage, Cake Filtration, And Membrane Morphology - Zydney
regions of the membrane that have first been covered by large aggregates. Just as importantly, this new theoretical framework provides the first quantitative analysis of the effects of the pore morphology on the rate of flux decline. Membranes with straight-through pores show the greatest rate of flux decline since the particles completely cover (block) the non-interconnected pores in these membranes. In contrast, particle deposition on the surface of a homogeneous membrane with highly interconnected pores will disturb the filtrate flow only over a relatively small penetration distance into the membrane pore structure. This implies that very thick membranes should exhibit a slower rate of flux decline due to the smaller relative disturbance in the flow, an effect that has been confirmed experimentally [ 16,17]. The flux decline for composite membrane structures shows a more complex behavior. The surface fouling completely blocks the pores in the upper skin layer, but the fluid rapidly redistributes itself through the very highly interconnected pores within the membrane substructure. This causes a shunting of the fluid flow towards the remaining open pores, leading to a reduction in the rate of flux decline compared to that for a single membrane layer. This theoretical analysis also provides a framework that can be used for the design and development of new membrane structures having reduced rates of fouling. This would include the proper choice of pore connectivity, membrane thickness, and the specific properties of the individual layers in more complex composite structures. These phenomena have often been neglected in membrane design/development, although they can clearly have a dramatic effect on the flow distribution, fouling, and overall selectivity for the particular separation. REFERENCES [1]
P. H. Hermans and H. L. Bred6e, "Principles of the Mathematical Treatment of Constant-Pressure Filtration," J. Soc. Chem. Intl. 55T (1936) 1.
[21
V. E. Gonsalves, "A Critical Investigation on the Viscose Filtration Process," Ree. Trav. Chim. Des Pays-Bas, 69 (1950) 873.
[3]
H. P. Grace, "Structure and Peformance of Filter Media," AIChE J., 2 (1956) 307.
[4]
J. Hermia, "Constant Pressure Blocking Filtration Laws - Application to Power Law Non-Newtonian Fluids," Trans. Inst. Chem. Eng., 60 (1982) 183.
[5]
K. Luckert, "Model Selection Based on Analysis of Residue Dispersion Using SolidLiquid Filtration as an Example," Int. Chem. Eng., 34 (1994) 213.
[6]
W. R. Bowen, J. I. Calvo, and A. Hemandez, "Steps of Membrane Blocking in Flux
Decline During Microfiltration," d. Membrane Sei., 101 (1995) 153. [7]
E. Iritani, Y. Mukai, Y. Tanaka, and T. Murase, "Flux Decline Behavior in Dead-End Microfiltration of Protein Solutions," J. Membrane Sci., 103 (1995) 181.
43
Fouling Phenomena During Microfiltration: Effects Of Pore Blockage, Cake Filtration, And Membrane Morphology - Zydney
[8]
J-Y. Wang, K-S. Chou, and C-J. Lee, "Dead-End Filtration of Solid Suspension in Polymer Fluid through an Active Kaolin Dynamic Membrane," Sep. Sei. Tech., 33 (1998) 2513.
[9]
S. T. Kelly, W. S. Opong, and A. L. Zydney, "The Influence of Protein Aggregates on the Fouling of Microfiltration Membranes During Stirred Cell Filtration," J. Membrane Sei., 80 (1993) 175.
[10]
W. Yuan and A. L. Zydney, "Humic Acid Fouling During Microfiltration," J. Membrane Sci., 157 (1999) 1.
[11]
G. Belfort, R. H. Davis, and A. L. Zydney, "The Behavior of Suspensions and Macromolecular Solutions in Crossflow Microfiltration," J. Membrane Sei., 96 (1994) 1.
[12]
C. C. Ho and A. L. Zydney, "A Combined Pore Blockage and Cake Filtration Model for Protein Fouling during Microfiltration." J. Coll. Interf. Sei. 232 (2000) 389.
[13]
C. C. Ho and A. L. Zydney, "Measurement of Membrane Pore Interconnectivity," J. Membrane Sci., 170 (2000) 101.
[141
W. Yuan, A. Kocic, and A. L. Zydney, "Analysis of Humic Acid Fouling using a Combined Pore Blockage - Cake Filtration Model," J. Membrane Sei. (in press 2002).
[15]
C. C. Ho and A. L. Zydney, "Protein Fouling of Asymmetric and Composite Microfiltration Membranes," Ind. Eng. Chem. Res., 40 (2001) 1412.
[16]
C. C. Ho and A. L. Zydney, "Effect of Membrane Morphology on Protein Fouling During Microfiltration," d. Membrane Sei., 155 (1999) 261.
[17]
C. C. Ho and A. L. Zydney, "Theoretical Analysis of the Effect of Membrane Morphology on Fouling during Microfiltration." Sep. gel. Teeh. 34 (1999) 2461.
[18]
L. Palacio, C. C. Ho, and A. L. Zydney, "Application of a Pore Blockage - Cake Filtration Model to Protein Fouling during Microfiltration," Bioteeh. Bioen~ (submitted 2002).
44
New Insights into Membrane Science and Technology: Polymeric and Biofunctional Membranes D. Bhattacharyya and D.A. Butterfield (Editors) © 2003 Elsevier Science B.V. All rights reserved.
Chapter 3
Differential scanning calorimetry and rheological experiments to study membrane formation via thermallyinduced phase-separation P.C. van der Heijden, M.H.V. Mulder, M. Wessling University of Twente, PO Box 217, 7500 AE Enschede, The Netherlands, Fax: xx31-53-4894611 1.
INTRODUCTION
1.1
Thermally-Induced Phase-Separation Porous polymer membranes can be prepared by different techniques: track-etching, stretching and phase-separation [ 1]. The phase-separation method is very common to prepare polymeric membranes and much attention has been paid to prepare membranes from various polymers and diluents [1,2]. The preparation of a porous structure consists of two steps. First, the homogeneous polymer solution has to undergo liquid-liquid demixing to obtain a polymer-rich continuous matrix and a dispersed polymer-lean phase of almost pure solvent. The second step is a fixation step to give the structure mechanical stability. Liquid-liquid demixing takes place when the solvent power is not sufficient anymore to keep the polymer dissolved. Two major methods can be distinguished to decrease the solvent power, either by adding a non-solvent which is called diffusion-induced phase-separation (DIPS) or by changing the temperature, defined as thermally-induced phase-separation (TIPS). The latter method is studied in this paper. Fig. 1 shows a schematic representation of the formation process of a porous structure with the TIPS method. The dashed arrow represents a typical cooling trajectory. When passing the binodal, the homogenous polymer solution liquid-liquid demixes in a polymer-rich phase with a polymer concentration given by the binodal and a polymer-lean phase of almost pure diluent. After liquid-liquid demixing of the binary polymer-diluent system, the structure has to be fixed. This can be achieved by crystallization, vitrification (glass transition), or gelation of the polymer-rich phase. After the structure fixation, it is possible to remove the polymer-lean phase by evaporation or extraction. An advantage of the TIPS method over the DIPS method is the ability to form porous structures with polymers which do not form a solution at room temperature, for example polyethylene and polypropylene.
45
Differential Scanning Calorimetry And Rheological Experiments To Study Membrane Formation Via Thermally-Induced Phase-Separation Van Der Heijden
,soo
c
H I binodal
. _ _
•
400 1
E(L) 300 -F-
"~pi~n°dalT
200
' 0.0
I 0.2
S
'
I 0.4
i
I 0.6
i
I 0.8
i 1.0
Volume fraction polymer (-) Figure 1. Example of a phase diagram of a binary polymer- diluent solution. The black arrow represents a cooling trajectory to obtain a porous structure with the TIPS method. It goes from the homogeneous solution (I-I) via the liquid-liquid demixing gap (L-L) into the structure fixation area (S). ~c is the critical concentration of the solution. The porous structures obtained with TIPS can be used for more applications than membranes, for example: separators m electrochemical cells, synthetic leather, 'breathable' rainwear, diapers, surgical dressings and bandages [3], inertial confmed fusion targets [4,5], biodegradable implants for cell transplantation [6], and scaffolds for tissue engineering [7]. Typical cell sizes obtained with this method are in the order of micrometers.
1.2
Historical Background
The formation of porous structures via TIPS was firstly reported by Castro [8]. In this patent, he investigated hundreds of polymer-diluent systems, both amorphous and crystalline polymers, on their ability to form a porous structure. Later on, Shipman [9] described a procedure to prepare porous structures with a certain shape, and Josefiak et al. [10] introduced a second liquid prior to the cooling step to adjust the morphology of the porous structure. The first papers on the TIPS method were published in 1984 [ 11 ]. Since then, the interest in the formation of porous structures of various polymerdiluent systems with the TIPS method has increased rapidly. Table 1 gives a
46
Differential Scanning Calorimetry And Rheological Experiments To Study Membrane Formation Via Thermally-Induced Phase-Separation Van Der Heijden
summary of crystalline polymer-diluent systems used in the TIPS method and Table 2 gives an overview of the formation of amorphous porous structures via TIPS. Besides the studies on new polymer-diluent systems, studies were performed on variations of the TIPS method by introducing, for example, an evaporation step to obtain an asymmetric structure [ 12-16].
Table 1. Crystalline polymer-diluent systems which have been used to form a porous structure with TIPS. Polymer Diluent Isotactic polypropylene (iPP) Diphenylether [ 17-19] n,n-Bis(2-hydroxylethyl)tallowamine [11,20-231 Polyvinylidene fluoride Cyclohexanone [ 11 ] (PVDF) Butrolactone Propyl enecarbonate Carbitolacetate Isotactic polystyrene (iPS) Nitrobenzene [24] Cyclohexane [25] Dioxane / isopropanol Poly(tetrafluorethylene-co-per- Chlorotrifluorethylene [26] fluoro-(propyl vinylether)) (Teflon®PFA, NeoflonT~FA) Poly(2,6-dimethyl- 1,4Cyclohexanol [ 12] phenylene ether) (PPE) Nylon 12 Poly(ethylene~lycol) [271 Dioxane / water [28] Poly(7-benzyl-l-glutamate) Benzene (PBIG) Dichloroethane Polyethylene (HDPE) Diphenylether [29] 1,2-Ditridecylphtalate / hexadecane[30] Polysilastyrene (PSS) Cyclohexane [31 ] Benzene 4-Methyl- 1-pentene (PMP) Diisopropylbenzene [5] Polylactide 1,4-Dioxane [6] As mentioned before, the formation of a porous structure with the TIPS method consists of two steps: liquid-liquid demixing to obtain a porous structure and the fixation step to obtain mechanical stability. Both steps have been studied separately for quite a time. The first papers on liquid-liquid demixing of polymer solutions were already published in 1950's (for references see [43,44]).
47
Differential Scanning Calorimetry And Rheological Experiments To Study Membrane Formation Via Thermally-Induced Phase-Separation Van Der Heijden
The existence of crystalline polymers has been known from the mid 1940's (a historic overview can be found m [45]) and the vitrification temperature depression (or glass transition temperature depression), for example, was already quantified in 1961 [46]. Table 2. Amorphous polymer-diluent systems which have been used to form a porous structure with TIPS. Polymer Diluent Atactic poly(methylmethacrylate) Cyclohexanol [32,33] (aPMMA) Tetramethylene sulfone (sulfolane) [34,35] 1-Butanol [32] Tert-butylalcohol [36] Atactic polystyrene (aPS) Cyclohexane [37] Diethylmalonate Cyclohexanol [38-41] Trans-decahy dronaphtalene (decalin) [42]
Liquid-liquid demixing of polymer solutions is extensively studied since the last half of the twentieth century (for references see [43,44]). Phase diagrams of many polymer-diluent systems have been determined visually or with other optical techniques such as optical microscopy and light scattering [21,23,27,28,35,38,43]. These techniques are also used to follow the time dependency of the demixing process to study the kinetics of demixing [17,27,33,47-50]. Also other experimental techniques have been occasionally used to compose phase diagrams or to follow the demixing process like viscometry [51 ], NMR [34], X-ray scattering [52] and DSC [32,42,53,54]. 1.3
Aim of this Paper In this paper, DSC is used to study liquid-liquid demixing and the structure fixation step of an amorphous polymer solution. In addition, rheological experiments are carried out to study liquid-liquid demixing as well. Using DSC and theology, new insights may be obtained in the membrane formation process. However, practical reasons can be mentioned to use these techniques instead of the frequently used light based techniques. For example, with light based techniques the refractive index between both the diluent and polymer has to differ significantly, and transparency of the experimental setup is required to observe a signal. This is not a problem with DSC and rheological experiments. Furthermore, with DSC and rheology it is possible to detect both the liquid-liquid demixing step and the structure fixation step in one experiment. 48
Differential Scanning Calorimetry And Rheological Experiments To Study Membrane Formation Via Thermally-Induced Phase-Separation Van Der Heijden
With optical microscopy for example, liquid-liquid demixing can be observed but it is impossible to observe the glass transition. Finally, high temperatures are involved in the TIPS process and to exclude evaporation of diluent, closed sample holders are necessary to use, and this is very easy to achieve with DSC. Q
DIFFERENTIAL SCANNING CALORIMETRY [55]
2.1
Introduction With DSC, the heat flow is measured as a function of time or temperature to observe physical transitions. After the introduction of the first differential scanning calorimetry (DSC) apparatus polymers have been studied extensively. See Refs. [56,57] for a historical overview about the development of thermal analysis devices in general and for polymers in particular. DSC is a well-known technique to study solid-liquid demixing (crystallization) [20,23,24,29] and vitrification [32,38,54] of liquid-liquid demixed polymer solutions. Berghmans and co-workers [32,42,53,54] used DSC as well to determine the liquid-liquid demixing temperature. They carried out DSC experiments for the systems atactic polystyrene / decalin and atactic polymethylmethacrylate / 1-butanol and cyclohexanol respectively. Upon cooling (cooling rate 5 K.min-1), an exothermic heat flow shift was observed and the onset was taken as the liquid-liquid demixing temperature. This signal agreed very well with optical observations. With one DSC run they could determine both the liquid-liquid demixing temperature and the glass transition temperature of the polymer-rich phase. However, DSC is hardly used for the determination of liquid-liquid demixing because the heat effect is very small and disappears easily in the base line drift. Recently, a rather new technique has been developed: temperature modulated differential scanning calorimetry (TMDSC) [58,59]. In spite of some discussion about the interpretation of the measured signals [60-67], TMDSC is a very useful device to measure small heat effects and to separate overlapping thermal events. In the next sections, TMDSC experiments are used to examine the cooling trajectory of binary polymer solutions in the concentrated region (at the right side of the critical point in Fig.l). The experimental work described in this work is carried out with the polymer-diluent system atactic polystyrene in 1dodecanol. A number of reasons can be mentioned to choose this system. This system forms a porous polymer structure with the TIPS method at room temperature [8], 1-dodecanol is a non-toxic solvent, polystyrene has solvents at room temperature and therefore easy to use, and many physical data are available for both components.
49
Differential Scanning Calorimetry And Rheological Experiments To Study Membrane Formation Via Thermally-Induced Phase-Separation Van Der Heijden
2.2
Temperature Modulated Differential Scanning Calorimetry
An extensive description of this technique can be found in Refs. [58,62,68]. Here, only a short description is given. With TMDSC, a second function (for example a sine wave) is superimposed onto the conventional linear or isothermal temperature ramps (see Fig. 2).
200
-
0 o v
"~ 1 9 5 O.
E I--
"
190 I
0
'
I
1
'
I
'
2
I
3
'
I
4
'
5
Time (min)
Figure 2: Linear cooling trajectory (thick black line) with superimposed temperature modulations.
The temperature ramp can then be described as: (1)
T = To + bt + A sin cot
where To is the initial temperature, b the underlying scanning rate, A the temperature amplitude, co the angular frequency, and t the time. The TAinstruments user manual [68] suggests a value of the underlying scanning rate (b) below 5 K.min "1 and typical values for A and co are 1 K and 60 s. The resulting heat flow consists of two contributions. The first part is caused by rapid processes and is proportional to the scanning rate. The second part is caused by kinetically hindered or irreversible processes and hence independent of the scanning rate: dQ dt
= Cp
dT -~
(2)
+ f (t,T) "
50
Differential Scanning Calorimetry And Rheological Experiments To Study Membrane Formation Via Thermally-Induced Phase-Separation Van Der Heijden
dQ/dt is the heat flow, Cp the heat capacity, and fit, T) a contribution to kinetic effects. The resulting heat flow is a wave function as well. With the help of discrete Fourier transformation software which resides in the DSC module, the heat flow signal is deconvoluted in a cyclic signal and an underlying signal (which is equivalent to the conventional DSC). In this work the modulus of the complex heat capacity is used to present the TMDSC results. This quantity is calculated from only the amplitude of both the temperature and the heat flow modulation. It is thus completely derived from the cyclic signal. Physical processes, which take place at time scales smaller or comparable to the modulation period, contribute to this complex heat capacity. Examples of such physical processes are vibrations and rotations of atoms. Slow physical processes, like the mobility of vitrified polymers with time scales much larger than a modulation period does not contribute to the complex heat capacity.
(Icp*l)
(IcjI)
2.3
Experimental
2.3.1 Materials Atactic polystyrene (aPS) (Styron* 686E) was kindly supplied by Dow Benelux NV (Mw and Mw/M,: 2.3.105 g.mol 1 and 2.1 respectively, determined with GPC). The diluent, 1-dodecanol (purity > 98%, Merck-Schuchardt), was used without further purification.
2.3.2 Sample preparation A homogeneous solution of aPS and 1-dodecanol was prepared in a threeneck bottle under nitrogen at 200°C. 1-Dodecanol vapor was allowed to evaporate during stirring with a mechanical stirrer. Small amounts of various polymer concentrations were poured in Petri-dishes and cooled in air. The compositions of the samples were determined by thermogravimetric analysis. About 20 mg of the sample was inserted in a platinum sample pan of a TGA 2950 Thermogravic Analyzer of TA Instruments and heated up to 200°C with a heating rate of 10 K-min-1. Afterwards the temperature was kept constant at 200°C for maximum 2 h to evaporate all the 1-dodecanol. From the ultimate weight loss the polymer concentration was determined.
2.3.3 Temperature (TMDSC)
Modulated Differential
Scanning Calorimetry
The TMDSC used is a DSC 2920 of TA Instruments. Calibration with indium and high density polyethylene (HDPE) (for calibration of the heat capacity) was carried out. About 5 mg of the sample was put in the aluminum closed sample pan. The TMDSC was heated to 200°C and kept isothermally for
51
Differential Scanning Calorimetry And Rheological Experiments To Study Membrane Formation Via Thermally-Induced Phase-Separation Van Der Heijden
30 min to ensure homogeneity. The cooling rate was set to 2 K-min 1 to 0°C and after an isothermal step of 5 min the sample was heated again with 2 K.min -1. The amplitude of the superimposed sine wave was 1 K with a period of 60 s [68]. The glass transition temperature Tg and the liquid-liquid demixing temperature TL_Las well as the heat capacity shift at TL_Lwas determined with the TA Universal Analysis software.
2.3.4 Optical Microscopy (OM) To compare the TMDSC results with a well-known technique for studying liquid-liquid demixing, optical microscopy experiments are carried out. The polymer sample was placed on an object glass within an aluminum ring (thickness 0.1 mm, inner diameter 5 mm) and covered by a second glass. To prevent diluent loss caused by evaporation, laboratory grease was used to stick the aluminum spacer to the object glasses [23]. The sample was placed in a hot stage (Linkam THMS 600) which was controlled by the Linkam TMS92 hot stage controller. The sample was heated and cooled with a rate of 2 K.min -1 and demixing was observed visually with an Olympus BH2 microscope (magnification 200x).
2.4
Results
Cooling and subsequent heating curves of aPS - 1-dodecanol are plotted in Fig. 3 for two polymer concentrations (weight fractions of 0.38 and 0.69). The modulus of the complex heat capacity (Icp'l) is plotted at the y-axis. Two transitions can be observed: the glass transition and a small baseline shift at higher temperatures, which is assumed to be the liquid-liquid demixing temperature. In the following, the onset of this signal upon cooling is defined as the liquid-liquid demixing temperature (TL-L) comparable with the observations of Amauts et al. and Vandeweerdt et al. with the conventional DSC [32,42,53,54]. The glass transition temperature (Tg) is chosen as the onset upon cooling because below this temperature influences can be expected of vitrification on the liquid-liquid demixing behavior. The difference in heat capacities ([Cp*l) between the results for the two polymer concentrations can be explained by the heat capacity difference between both single components. Literature values of the heat capacity for pure polystyrene and 1-dodecanol at T = 180°C are 2.1 J.gl.K-1 [69] and 3.1 J.gl.K-1 [70], respectively. Therefore, a polymer solution that contains the higher content of polystyrene should have a smaller heat capacity than a diluted polymer solution. At least, when it is assumed that the heat capacity of a solution of aPS in 1-dodecanol is between the values of both pure components.
52
Differential Scanning Calorimetry And Rheological Experiments To Study Membrane Formation Via Thermally-Induced Phase-Separation Van Der Heijden
3.0
'7, '7
--
2.5-ETJ
-..3
"
0
0.
2.0
--
..,.=,_.-
69 wt-%
1.5
;
I
;
40
I
80
:T,, ;
I'
120
'
I
160
'
200
Temperature (°C) Figure 3. Cooling and subsequent heating curves (weight fractions of polymer: 0.38 and 0.69). Gray lines" heating curves, black lines" cooling curves.
With conventional DSC experiments, an optimum has to be found between a high scanning rate which results in a high intensity and a low scanning rate which results in a high resolution of a thermal signal. Consequently, a small thermal transition such as liquid-liquid demixing is very difficult to observe with conventional DSC. With TMDSC, both a low scanning rate (the underlying scanning rate) and higher scanning rates (the temperature modulations) are present within one experiment and therefore it is very suitable to detect liquid-liquid demixing. In the cooling curves the L-L phase transitions at TL_Lare represented by a steep heat capacity shift. The heating curves have the same slopes as the cooling curves only at the liquid-liquid demixing temperatures the transition is not as distinct. This difference is discussed later in this section. In the following, the details of the cooling curves are further used and discussed only. Performing such TMDSC cooling experiments over a large concentration range allows the construction of the phase diagram of the polymer-diluent system. To support the assumption of the base line shift to stem from the L-L demixing, the TMDSC results are compared with optical microscopy (OM) data indicating visually the phenomenon of L-L demixing. In Fig. 4, the TL-Land Tg determined with TMDSC and OM are plotted. The open circles represent the TMDSC liquid-liquid demixing data whereas the filled black squares are the OM data. The closed circles are the glass-transition temperature data points. The liquid-liquid demixing data obtained from TMDSC
53
Differeraial Scanning Calorimetry And Rheological Experiments To Study Membrane Formation Via Thermally-Induced Phase-Separation Van Der Heijden
and OM agree well and therefore it can be concluded that the observed TMDSC signals are indeed caused by liquid-liquid demixing. 200
~,
160
o o v
::}
120
(1.) (:}.
E
80
!--
40
0.0
0.2 Weight
0.4 fraction
0.6
0.8
polymer
1.0 (-)
Figure 4. Phase diagram a P S - 1-dodecanol. Open circles: TMDSC data L-L demixing. Closed squares: OM data L-L demixing (cloud points). Closed circles: TMDSC data glass transition. Lines are drawn to guide the eye.
The shift of the TMDSC curve at the liquid-liquid temperature is assumed to be completely caused by the enthalpy of demixing. The difference in values in the heat capacity shift (Acp* at Tt.t, defined in Fig. 3) between the different concentrations is caused by the interaction between polymer and diluent. This can be quantified by calculating the enthalpy of mixing with the help of the Flory-Huggins theory [54]. To minimize the error in the calculation of the modulus of the complex heat capacity with the TMDSC software, it is recommended that at least 4 complete superimposed cycles fit within a phase transition [68]. This requirement is satisfied for the glass transition because this transition covers a temperature range of at least 10 K. However, in case of liquid-liquid demixing the heat capacity shift only covers a temperature interval of 2 K, so only one modulated cycle fits within this transition. By lowering the underlying cooling rate, the number of cycles within the transition can be increased; the resulting TMDSC curves are shown in Fig. 5. From Fig. 5 it can be concluded that cooling rates of 2 K.min -1 and lower have no significant influence on the measured modulus of the complex heat capacity. As already mentioned the measured complex heat capacity is only measured from the amplitudes of the modulated heat flow and modulated temperature, hence it is not influenced by the cooling rate. However, the shape of the curve can be influenced by time-dependent effects like the growth of the
54
Differential Scanning Calorimetry And Rheological Experiments To Study Membrane Formation Via Thermally-Induced Phase-Separation Van Der Heijden
demixing domains and the increasing viscosity of the polymer solution (see section 3) upon cooling. The growth of the demixed domains can be the reason for the difference between the cooling and heating curve (see Fig. 3 for a 38 wt.% polymer solution). In the demixed polymer solution regions of almost pure diluent grow when the temperature of the solution is in between the TL-L and the Tg. The TMDSC heating curve was measured after the cooling trajectory, so, the demixed domains already had a long time to grow. Therefore, it is more difficult for the polymer solution to follow the temperature modulation because of larger distances the diluent molecules must diffuse. ,,
,
,,
0.2 K.min~ 2 . 8
,,
,
,
2 K-min"~
-"
"7 "7
2.4 - -
v
-1
•
..
(3 J
2.0--
~/
5 K.min
1.6 J
50
!
I 100
'
I
'
150
I 200
Temperature (°C) Figure 5. Influence of cooling rate on modulus of the complex heat capacity. Weight fraction of polymer is 0.48.
The shape of the curves at low cooling rates (between 0.2 K.min 1 and 2 K.min1) are comparable, this means that influence of time depending effects play a negligible role in the cooling curves at these cooling rates. 2.5
Conclusions and Outlook With Temperature Modulated DSC, liquid-liquid demixing of polymerdiluent systems can be determined as well as the glass transition of the polymerrich phase in one run at a relatively low scanning rate. Liquid-liquid demixing observed with TMDSC agrees well with visually observed cloud points.
55
Differential Scanning Calorimetry And Rheological Experiments To Study Membrane Formation Via Thermally-Induced Phase-Separation Van Der Heijden
Upon quantifying the heat capacity shift at the liquid-liquid demixing temperature additional information with respect to the thermodynamics of the polymer solutions can be obtained. Furthermore, it offers the possibility to follow the heat capacity shift as a function of temperature both theoretically and experimentally to study the kinetics of liquid-liquid demixing. Especially this is interesting in the region of the glass transition temperature in which the structure fixation step can be studied. 3.
VISCOMETRY AND RHEOLOGY
3.1
Introduction
During the formation of porous structures with the TIPS process, the homogenous solution is demixed in two liquid phases and afterwards the polymer-rich phase is vitrified. It is expected that these phase transitions can be observed with rheological experiments. The possibility of studying liquid-liquid demixing with viscometry and rheology experiments is briefly discussed in this section. The experiments are carried out for the polymer-diluent system aPS diisodecylphthalate (DIDP). The reason to study this system is the following. The high vapor pressure of many common diluents in relation to the liquidliquid demixing temperature is often a problem studying the thermally-induced phase-separation method experimentally. Special care has to be taken to prevent evaporation of the diluent during an experiment. The diluent DIDP shows a cloud point with the polymer of about 50°C and has a very low vapor pressure. (For comparison, the vapor pressure of water at T = 20°C is 2.10 3 Pa and the vapor pressure of DIDP at T = 100°C is 0.1 Pa [70].) Therefore, for this system no special care has to be taken to avoid evaporation of diluent.
3.2
Experimental Homogeneous solutions of aPS in DIDP (Merck-Schuchardt, purity>99%) were prepared in the same way as described in section 2.3.2. The viscosity of the polymer diluent system aPS-DIDP was measured with a Brabender Viscotron, Mod. Nr 8024. Starting at high temperatures (in the homogeneous solution), the viscosity was measured for different shear rates at a fixed temperature. Subsequently the temperature was lowered and the complete procedure was repeated. Furthermore, polymer solutions of aPS in DIDP were studied in a Bohlin Rheometer CS50. Upon cooling with a cooling rate of 2 K.min -1, the complex viscosity was measured for different strains at a constant frequency of 0.1 Hz.
56
Differential Scanning Calorimetry And Rheological Experiments To Study Membrane Formation Via Thermally-Induced Phase-Separation Van Der Heijden
3.3
Results 200
150 --t~ ¢~
13_ >"100 0 ¢o
>
50-
~'~1 I 20
i
40
I 60
s~ = 80
T e m p e r a t u r e (°C) Figure 6. Viscosity of a 30 wt.% solution of aPS in DIDP as a function of the temperature upon cooling for different shear rates (indicated in the figure). Results of the viscosity measurements are visualized in Fig. 6 for a polymer concentration of 30 wt.%. Experiments in the concentration range of polymer between 20 and 40 wt.% show a comparable trend'as plotted in Fig. 6. The observation of an increasing viscosity upon cooling as the temperature approaches the liquid-liquid demixing temperature is in agreement with the work of Wolf et al. [51]. In fact, the location of the sharp decrease in viscosity (between T = 45 and 50°C) is in good agreement with cloud point experiments of this polymer-diluent system (T = 49°C). The observation of the sudden decrease at lower temperatures inside the liquid-liquid demixing gap is not in agreement with the expectation. Upon cooling a polymer solution, the viscosity increases. At the liquid-liquid demixing temperature, a polymer-rich matrix is formed enclosing the polymerlean phase. The viscosity is expected to be dominated by the viscosity of the continuous, polymer-rich, phase of the system. Therefore, it should be expected that upon liquid-liquid demixing the viscosity increases further. An explanation of the sharp decrease in viscosity could be the loss of contact between the polymer rich matrix and the rotating cone of the viscometer. Upon storage below the cloud point temperature the demixed polymer solution displays synereses; DIDP is expelled out of the continuous phase. This process is probably enhanced by the applied shear rate resulting into the formation of a slip
57
Differential Scanning Calorimetry And Rheological Experiments To Study Membrane Formation Via Thermally-Induced Phase-Separation Van I)er Heijden
layer of DIDP between the polymer rich matrix and the rotating cone. To apply less mechanical force on the polymer solution, an oscillating rheometer was used instead of the rotating viscometer. 2 D1
?
t
Q... >,
.~_ f/)
o o
U)
•~
0.5
•
x
1
Q.,
E o
20
30
40
50
60
70
Temperature (*C) Figure 7. Complex viscosity for a 30 wt.% solution of aPS in DIDP. The frequency is 0.1 Hz. The strain varies from 0.1 to 0.001. The cooling rate is 2 K-min-1. In Fig. 7, a small plateau can be observed upon cooling at the liquid-liquid demixing temperature in the complex viscosity at a low strain (0.001). However, as expected, the complex viscosity increases upon further cooling. With increasing strains the plateau becomes longer and the complex viscosity even shows a depression at a strain of 0.1. The reason of this depression is probably the same as in the viscosity experiments. Diluent is pushed out of the demixed solution and forms a slip layer. When using rheology or viscometry experiments to determine the liquidliquid demixing temperature, sufficient mechanical force should be applied to achieve immediate synereses of the demixed solution. The synereses results into an apparent viscosity drop that could be used as a marker for the onset of demixing. However, for studying the mechanical properties of the demixed solution, very small strains are recommended in order to avoid shear enhanced synereses. Otherwise, the measurement of the mechanical properties of the demixed solution is unreliable.
58
Differential Scanning Calorimetry And Rheological Experiments To Study Membrane Formation Via Thermally-Induced Phase-Separation Van Der Heijden
3.4
Conclusions and Outlook Liquid-liquid demixing temperatures can be observed with rheology and viscometry by using high shear rates. To study the mechanical behavior of demixing polymer solutions, care has to be taken to avoid altering of the liquid structure due to shear forces. It should be of much interest when rheological experiments can be carried out with a polymer solution at different frequencies. These results can be compared with results of emulsions to obtain information about the structure belonging to the demixed solution. Furthermore, rheological experiments can be very suitable to study the mechanical behavior of the polymer-diluent system at a temperature in the region of the structure fixation step. For example, the system aPS - DIDP forms a gel at room temperature because of the high viscosity of the polymer solution. After the gelation of the solution, no growth of the demixed domains is observed. Hence, gelation is sufficient to stop the liquid-liquid demixing process. The polymer-diluent system aPS - 1-dodecanol, which was studied in section 2, shows an onset of the glass transition temperature at about 65°C. The glass transition temperature is known to be sufficient to stop the liquid-liquid demixing process. However, it is very interesting to study whether at higher temperatures than the glass transition temperature, the liquid-liquid demixing process is influenced by or even stopped by gelation. Rheological experiments have the potential to clarify the structure fixation step of a polymer solution.
4.
ACKNOWLEDGEMENTS
E. Schomaker, J. Bos en R. Lammers of Akzo Nobel in Amhem are acknowledged for discussions with respect to the TMDSC work, and for giving us the possibility to carry out the TMDSC experiments. B. Reuvers (Akzo Nobel in Arnhem) and M. van Egmond (University of Twente) are acknowledged for carrying out the rheological experiments and for the discussions. REFERENCES
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Differential Scanning Calorimetry And Rheological Experiments To Study Membrane Formation Via Thermally-Induced Phase-Separation Van Der Heijden
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Non-Invasive Characterizations of Membrane Fluid Transport and Fouling
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New Insights into Membrane Science and Technology: Polymeric and Biofunctional Membranes
D. Bhattacharyya and D.A. Butterfield (Editors) © 2003 Elsevier Science B.V. All rights reserved.
Chapter 4
Study of membrane fouling and cleaning in spiral wound modules using ultrasonic time-domain retlectometry Zh.-X. Zhang a, A.R. Greenberg b, W.B. Krantz ¢, and G.-Y. Chai" aDepartment of Chemical Engineering, University of Colorado at Boulder, Boulder CO 80309-0424, USA bDepartment of Mechanical Engineering, University of Colorado at Boulder, Boulder CO 80309-0427, USA CDepartment of Chemical Engineering, University of Cincinnati, Cincinnati, Ohio 45221-0171, USA NSF Membrane Applied Science and Technology Center, University of Colorado, Boulder, CO 80309-0424, USA ABSTRACT Better detection and control of fouling in liquid-separation processes is essential if membranes are to find increased use in a variety of industrial applications. Ultrasonic time-domain reflectometry (UTDR) is an in situ, noninvasive real-time technique that has been successfully utilized to quantify fouling and cleaning of fiat sheet membranes. This study describes the extension of UTDR for the measurement of fouling and cleaning in commercial membrane modules employing spiral-wound elements. Experiments were conducted using a high-pressure separation system incorporating a commercial spiral-wound module. Calcium sulfate was utilized as the primary scalant in concentrations ranging from 1.2-1.8 g/L; ferric hydroxide colloidal fouling also occurred during the tests. A multi-cycle protocol was adopted whereby each cycle consisted of three stages: pure water equilibration, fouling and cleaning. Permeation rate as well as acoustic amplitude and arrival time measurements were made at regular intervals during three complete cycles on two modules. A new signal analysis protocol was developed such that systematic changes in the entire acoustic spectrum as a function of module operation time could be represented in terms of shift factors. In addition, gravimetric, microscopic and spectroscopic measurements were made at the end of the experiments so that the extent of the membrane fouling and cleaning could independently assessed. Results indicate that the overall decline in permeation rate is reasonably correlated with changes
65
Study Of Membrane Fouling And Cleaning In Spiral Wound Modules Using Ultrasonic Time-Domain Reflectometry - Greenberg
in both the amplitude and arrival time shift factors. Overall, this work confirms the effectiveness of the UTDR technique and provides a sound basis for continued development with an ultimate goal of providing on-line sensors for the timely detection of fouling and cleaning of commercial liquid-separation membrane systems. 1.
INTRODUCTION
As generally def'med, membrane fouling involves the deposition of retained particulates, colloids, macromolecules, crystallization of dissolved inorganic salts on or within the membrane. Fouling is the most critical problem limiting further growth and wider application of membrane-based liquidseparation processes [ 1]. Although the occurrence of fouling can be traditionally inferred from a marked decline in permeate flux or product quality, such an inference is unreliable since other phenomena such as concentration polarization, membrane compaction and degradation may also lead to similar changes in permeate flux or quality [2]. Moreover, using currently available techniques, it is not possible to quantify when a commercial membrane module has been adequately cleaned via chemical or mechanical de-fouling. Consequently, there is a compelling rationale for the development of an in situ, noninvasive technique that can monitor the growth and removal of a membrane-fouling layer in real time under realistic operating conditions. Such a technique would not only provide for identification of the specific phenomena underlying the flux decline, but also enable the improvement of fouling remediation strategies. Ultrasonic time-domain reflectometry (UTDR) has been used in a number of applications that involve the measurement of the location and topographical characteristics of a moving or stationary interface [3]. When the operating requirements of commercial high-pressure separation modules are considered, UTDR appears to be more suitable for real-time monitoring of membrane fouling than other noninvasive techniques, such as nuclear magnetic resonance microimaging [3]. Indeed, previous studies have demonstrated that UTDR can be successfully utilized to quantify fouling in a flat-sheet membrane module [2-4]. Recently, the UTDR technique was extended to the much more complex situation that occurs with the multiple layers that comprise spiral-wound reverse-osmosis (RO) modules. As in the flat-sheet studies, this work utilized calcium sulfate as the primary scalant due to its importance as a membrane-fouling agent [5-8]. Whereas the initial efforts to interpret the complex acoustic spectra utilized only a few of the many signal peaks [9,10], the present study employed a signal
66
Study Of Membrane Fouling And Cleaning In Spiral Wound Modules Using Ultrasonic Time-Domain Refleetometry - Greenberg
analysis protocol that accounts for systematic shifts in the entire acoustic spectrum as a function of module operating time. This paper also reports the first results obtained for UTDR detection of deposits formed by multiple foulants. 0
2.1
ULTRASONIC TIME-DOMAIN REFLECTOMETRY (UTDR)
Basic Concepts
The basic principles of UTDR and their application to the measurement of fouling on a fiat-sheet membrane have previously been described [2,3]. As shown in Figure 1, the ultrasonic wave pulses are transmitted by the transducer and propagate through the module. When the pulses encounter an interface, such as the one formed by the feed solution and top surface of the membrane, they will be reflected back to the transducer. The amplitude of the reflected waves depends on the acoustic impedance difference between the media on either side of the interface and the topography of the interface. The acoustic impedance is a function of the physical characteristics of the medium, and is given by the product of the density and the acoustic velocity within the medium. In addition, the time required for the transmission and reflection of a wave from the interface (i.e. arrival time) is a function of the path length and wave velocity: 2As AT = ~
(1)
where AT is the arrival time, As is the distance between the transducer and the interface and c is the acoustic velocity in the medium. Since the impedance, interface properties and path length may change with the growth of a fouling layer on the membrane surface, the change in the amplitude (AV) and the shift in the arrival time (AT) of the interface echoes can be analyzed and used to monitor fouling quantitatively in real-time.
2.2
Adaptation for Spiral-Wound Membrane Modules When the UTDR technique is applied to a spiral-wound RO membrane module contained in a cylindrical fiber-reinforced polymer housing, the signal analysis is complicated by two factors. First, a much more complex signal pattern will be obtained due to multiple reflections from the outer and inner surfaces of each layer of the multiple windings in the module assembly (Figure 2) [ 10]. Second, a significant loss of acoustic information will be caused by signal attenuation through these layers as well as through the module shell. The various reflections from these interfaces correspond to the many peaks in the waveform
67
Figure 1. Schematics that represent the principles of UTDR measurement: (Left) cross-sectional view of the separation cell, the externally mounted ultrasonic transducer, and the primary reflections, 1, 2, and 3 from the cell plate/fluid, fluidmembrane, and fluidfouling layer interfaces, respectively; (Right) the general ultrasonic response in the time-domain for each of these reflections.
Study Of Membrane Fouling And Cleaning In Spiral Wound Modules Using Ultrasonic Time-Domain Reflectometry - Greenberg
obtained by plotting the signal amplitude as a function of the arrival time (Figure 3). Here the peaks in the waveform represent very complex superposed reflections from the many close-packed interfaces within the spiral-wound module instead of the simpler patterns from the largely separated interfaces that are characteristic of the fiat-sheet geometry. Within this complex spectrum, certain peaks are more responsive to the growth and subsequent removal of fouling layers than others. In addition, the shape of the waveforms depends on the position and mounting mode of the transducer on the external surface of the module shell. In the previous studies [10], considerable effort was required to select and analyze those peaks in the amplitude versus arrival time spectra that appeared most sensitive to fouling. The technique employed was time-intensive and made real time analysis difficult. In addition, all of the acoustic information obtained could not be fully used. In order to overcome these limitations, a more sophisticated signal acquisition and analysis protocol was developed. This procedure accounts for systematic shifts in the entire acoustic spectrum as a function of module operating time and enables information about the extent of fouling to be obtained in real time. The specific steps of the signal analysis protocol are summarized below: • All acoustic spectra are digitized. • A reference spectrum is obtained for a particular set of operating conditions with the spiral-wound module of interest. The reference spectrum corresponds to a "zero" operating time at which fouling has not occurred. Changes in the acoustic spectra as a function of operating time are determined based upon a systematic departure or shift from the reference spectrum. Such changes are shown in Figure 3 for the RO module employed in this study. Here, two waveforms corresponding to the reference spectrum and an updated spectrum, i.e. a spectrum obtained at a later time, are superposed to demonstrate the changes in the signal amplitude and arrival time. • Two shift factors are computed using a root-mean-square calculation such that values are always positive. One shift factor, QA, is based upon changes in the signal amplitude and the other factor, QT, is based upon changes in the signal arrival time. The two relationships are: N
"QT = '
~(Tiu_Tir) 2 1/2 N
(2)
Here, Viu and Vir are the amplitude of peak i in the updated spectrum and the reference spectrum, respectively; Tiu and Tir are the arrival time of peak i in the updated spectrum and the reference spectrum respectively; and N is the total number of peaks in the spectrum under consideration.
69
Study Of Membrane Fouling And Cleaning In Spiral Wound Modules Using Ultrasonic Time-Domain Reflectometry - Greenberg
Transducer
Shell Feed Channel Spacer Membrane Permeate Collection Film
II II II [Layer1
~l ~]
Layern
-~ [
[
Figure 2. Schematic showing a cross-sectional view of a spiral-wound RO module assembly, including the fiberglass shell, the position of an externally mounted ultrasonic transducer, and reflections from the multiple interfaces. The width and length of the arrows are proportional to the reflected amplitude and arrival time, respectively.
2.50E-02
;~ 0.00E+00 3.7:E-0
J
E-05
t
AT (s)
°~,,i
-2.50E-02
Figure 3. Representative results showing a portion of the acoustic response in terms of signal amplitude versus arrival time (AT) from a commercial spiral-wound RO module. Changes in the UTDR signal peaks that occur as function of operating time are observed as a shift from the reference spectrum (heavy line) to the updated spectrum (light line).
The effectiveness of the new methodology was assessed by quantifying the signal fluctuations obtained during 30-hr runs using pure-water, and then representing
70
Study Of Membrane Fouling And Cleaning In Spiral Wound Modules Using Ultrasonic Time-Domain Reflectometry - Greenberg
the effects of module fouling in terms of the shift factors based upon anticipated changes in the signal amplitude and arrival time. An "ideal" response is shown in Figure 4 and indicates the expected relationship between the permeation rate and the amplitude shift factors. Here, the shift factors and the permeation rate are plotted as a function of operating time for a pure water equilibration step, a period of CaSO4 fouling, and a pure-water cleaning step. The signal analysis protocol is designed to distinguish the relatively small-scale fluctuations due to normal operating condition variations from the large-scale changes that are due to fouling. The dashed lines represent a "small-scale fluctuation" confidence interval for each parameter. From a practical point of view, a successful technique should indicate when the signal has moved out of the normal operating window, i.e. the confidence interval, as a result of module fouling. The new methodology also allowed instantaneous calculation of the shift factors so that the occurrence and removal of fouling in the RO spiral-wound module could be detected and monitored quantitatively in real-time. Moreover, the portion of the spectrum that provides the basis for the shift factor calculations can be easily altered. For example, Figure 5 shows a representative waveform in which the spectrum is divided into three regions, each designated by a different rectangle. The waveform within each rectangle contains information from different layers within the module. Since fouling does not occur to the same extent in all regions of the module at the same time, particular portions of the spectrum will be most responsive to its occurrence and thus provide the most sensitive detection of the growth and removal of fouling layers. 3. EXPERIMENTAL 3.1. RO System and UTDR Instrumentation Based on previous work [9,10], the experiments were conducted on a bench-scale RO flow system that allowed for the control and measurement of important system operating parameters including temperature, pressure, flow rate and calcium-sulfate feed concentration. The major components of the RO system consisted of a Koch 2521 spiral-wound RO membrane module, a feed tank with temperature control, a low-pressure pump used to transport the feed through a 2micron prefilter, and a high-pressure booster pump to drive the RO membrane separation process (Figure 6). The system was designed to enable both the retentate and permeate from the membrane module to be returned to the feed tank to enable stable operation and conservation of chemicals. The acoustic hardware included two specially designed high-frequency (3.5 MHz) transducers (Research Institute of Acoustics, Chinese Academy of Science), a pulsar-receiver (Panametrics 5052 PRX) and a digital oscilloscope (Nicolet Pro 50) that enabled a 20-ns sampling rate. A custom mounting assembly was utilized to allow the transducers to be securely attached at various
71
Operating Time (hr) Figure 4. Schematic showing the anticipated change in the UTDR shift factors (QA and QT)and the permeation rate using the new signal analysis protocol during the water equilibration (WE), fouling (F) and cleaning (C) phases of module operation. The solid lines represent the overall shift factor and permeation rate behavior as determined from values obtained at discrete time intervals (solid circles). The dashed lines indicate a confidence interval for normal operation. Fluctuations within the confidence interval are acceptable but departures outside the dotted lines are taken to indicate module fouling.
Study Of Membrane Fouling And Cleaning In Spiral Wound Modules Using Ultrasonic Time-Domain Reflectometry - Greenberg
positions on the external surface of the module. Secure attachment is essential to maximize changes in the acoustic waveforms that correspond to membrane fouling and cleaning. In the current experiments the transducers were mounted on the midstream and downstream portion of the module. Such multiple-point sampling can provide important information about the relative uniformity of the fouling and cleaning processes as a function of position. The UTDR signals were continuously monitored and recorded (waveform records containing 1000 data points) using a data-acquisition system that consisted of a personal computer (Pentium II 300 MHz) with a GPIB board (CyberResearch). Custom dataacquisition software enabled interfacing with the measurement instrumentation, real-time analysis of the UTDR signals, calculation of the two shift factors, and graphical display of both the shifts in the acoustic spectrum and changes in the values of the two shift factors. The acoustic hardware included two specially designed high-frequency (3.5 MHz) transducers (Research Institute of Acoustics, Chinese Academy of Science), a pulsar-receiver (Panametrics 5052 PRX) and a digital oscilloscope (Nicolet Pro 50) that enabled a 20-ns sampling rate. A custom mounting assembly was utilized to allow the transducers to be securely attached at various positions on the external surface of the module. Secure attachment is essential to maximize changes in the acoustic waveforms that correspond to membrane fouling and cleaning. In the current experiments the transducers were mounted on the midstream and downstream portion of the module. Such multiple-point sampling can provide important information about the relative uniformity of the fouling and cleaning processes as a function of position. The UTDR signals were continuously monitored and recorded (waveform records containing 1000 data points) using a data-acquisition system that consisted of a personal computer (Pentium II 300 MHz) with a GPIB board (CyberResearch). Custom dataacquisition software enabled interfacing with the measurement instrumentation, real-time analysis of the UTDR signals, calculation of the two shift factors, and graphical display of both the shifts in the acoustic spectrum and changes in the values of the two shift factors.
3.2. Experimental Procedure As indicated in Table 1, the testing protocol utilized three cycles, each of which generally contained water equilibration, fouling and cleaning phases. During the water-equilibration phase (phase A), the membrane module was operated using pure deionized water until the membrane module performance (permeation rate) and the UTDR signals (amplitude and arrival time) reached steady-state values. Then the feed was switched to an aqueous calcium sulfate solution that was prepared by adding anhydrous calcium sulfate powder (Aldrich, 325 mesh) to deionized water (equilibrium calcium sulfate solubility: 2 g/L).
73
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Study Of Membrane Fouling And Cleaning In Spiral Wound Modules Using Ultrasonic Time-Domain Refleetometry - Greenberg
This initiated phase B in which concentration polarization and fouling occurred. This phase was generally allowed to continue until the ultrasonic responsed permeation rate approached steady-state behavior. In the subsequent cleaning phase (phase C), the feed solution was switched back to pure deionized water in order to remove the fouling layer from the membrane surfaces. Overall, long-term experiments were conducted on two modules in order to accommodate the multiple fouling and cleaning cycles; the modules were run for a total of 654 and 1108 hours, respectively. In each experiment the fouling cycles generally utilized systematically higher concentrations of calcium sulfate. Table 1. Details of the Fouling and Cleaning Protocol
Module Cycle/Phase I Cycle 1:CaSO4 (1.2 g/L) Phase A: Water Equilibration Phase B: Fouling Phase C: Cleaning Cycle 2:CaSO4 (1.6 g/L) Phase A: Water Equilibration Phase B: Fouling Phase C: Cleaning Cycle 3:CaSO4 (1.6 g/L) Phase A: Water Equilibration Phase B: Fouling Total Operating Time (hr) II Cycle 1:CaSO4 (1.2 g~) Phase A: Water Equilibration Phase B: Fouling Phase C: Cleaning Cycle 2:CaSO4 (1.6 g/L) Phase A: Water Equilibration Phase B: Fouling Phase C: Cleaning Cycle 3:CaSO4 (1.8 g/L) Phase A: Water Equilibration Phase B: Fouling Phase C: Cleaning Total Operating Time (hr)
Duration (hr) 37 164 82 30 100 71 70 100 654 72 102 150 75 126 110 33 296 144 1108
All experiments were conducted under the same operating conditions. The upstream pressure of the module was maintained as 0.7MPa, and the feed
76
Study Of Membrane Fouling And Cleaning In Spiral Wound Modules Using Ultrasonic Time-Domain Reflectometry - Greenberg
solution flow rate was maintained as 0.9 L/min. The temperature of the feedsolution in the tank was controlled at 20°C and the inlet temperature of the module was maintained at 22°C. The CaSO4 concentration of the feed, permeate and retentate were measured at regular intervals using a conductivity meter (Yellow Spring Instruments), and the membrane rejection was calculated via standard methods. In addition, the permeate and retentate flow rates were determined via a digital balance and timer. In order to confirm independently the extent of membrane fouling indicated by the values of the acoustic shift, each of the two RO modules was opened after the appropriate phase in the operating protocol for visual inspection, morphological studies, and gravimetric and elemental analysis. Module I was opened at the end of the Cycle 3 fouling phase to observe the extent of coverage as well as the distribution of the fouling layer on the membrane surfaces, while Module II was opened at the end of Cycle 3 cleaning phase to determine the degree of fouling deposit removal via pure water cleaning (Table 1). The morphological features of the fouled and water-cleaned membrane samples were characterized using light microscopy (Nikon Labophot) and scanning electron microscopy (Aspex Instruments). Gravimetric analysis was performed o n sections of the unrolled membrane using a microbalance, and the elemental composition and distribution of the deposits were determined via energy dispersive spectroscopy (EDS). 4. RESULTS AND DISCUSSION The important aspects of the UTDR response are demonstrated in the data obtained for module I after the completion of each of the three cycles. Results for the first cycle are shown in Figure 7 for the two shift factors, QA and QT, from the downstream transducer as well as the permeation rate as a function of module operating time. In order to compare directly the sensitivity and effectiveness of the UTDR technique with the new signal analysis protocol and the traditional fouling detection mode, changes in both the shift factors and the permeation rate are represented as percentage changes from their respective reference values, i.e. the values at the start of phase A. During the water-equilibration phase, the permeation rate is relatively constant, whereas there is a gradual but distinct increase in QA. During the fouling phase, the permeation rate declines by approximately 12% and QA increases by over 1500%. By completion of the cleaning phase, the permeation rate has returned to its reference value; however, QA remains almost unchanged. In addition, no apparent change in QT is observed over the course of cycle 1. Note that the spikes in QT at the end of the fouling run and the beginning of the cleaning run are believed to be due to a problem with the power supply. Similar results for the shift factors were obtained from the midstream transducer. Overall, the observed behavior in QA and QT was not as expected. Nevertheless, we continued with the protocol indicated in Table 1.
77
Study Of Membrane Fouling And Cleaning In Spiral Wound Modules Using Ultrasonic Time-Domain Reflectometry - Greenberg
Results for the second cycle are shown in Figure 8. Here each of the shift factors is represented by three different curves, each of which corresponds to the different regions of the signal spectrum shown in Figure 5. These data demonstrate that the shift factors from each region of the spectrum lead to a similar response although the differences suggest that particular regions of the spectrum may indeed be more sensitive to the presence and removal of a fouling layer. After establishing steady-state behavior during the equilibration phase, there is a marked decline in the permeation rate (~20%) and an increase in QA and QT during the fouling phase. During the subsequent cleaning phase, the permeation rate and QT return to their reference values. Although QA clearly trends in the same direction, this factor has not reached the reference value at the completion of the cleaning cycle. The results for cycle 2 obtained from the midstream transducer, however, were considerably different than those from the downstream transducer. Indeed, the signals from midstream transducer provided shift factor responses that were rather similar to those observed during cycle 1 (Figure 7). This emphasizes a fundamental difference between the permeation and ultrasonic measurements whereby the former represents an average response for the entire module while the shift factors can reflect the condition of the module on a local basis. A third operating cycle was then conducted under the same conditions as for cycle 2. However, in this case the protocol was modified so that instead of a cleaning phase, the membrane module would be opened at the end of the fouling phase so that microscopic, spectroscopic and gravimetric analyses could be performed. Results for the permeation rate and shift factor responses during cycle 3 are shown in Figure 9. As expected, the three response parameters are reasonably constant during the water equilibration phase. With the initiation of the fouling phase, there is a relatively rapid decrease in the permeation rate (~20%) over the first 10 hours to an approximately steady-state value that is maintained for the next 90 hours until the end of the run. In contrast, both shift factors increase throughout the duration of the fouling phase. Whereas QA increases at a relatively constant rate, the initial rate of increase of QT becomes more modest after the first 30 hours of fouling. Overall, these results confirm the trends observed during cycle 2 (Figure 8). At the end of the fouling phase in cycle 3, the module was opened and unwound. Visual examination indicated the presence of a reddish-brown deposit on the membrane surface, with the color turning progressively lighter (whiter) from the entrance to the exit end of the module. Microscopic examination confirmed the coexistence of the reddish-brown and white deposits where the latter evidenced the same characteristics previously reported for crystalline CaSO4 (Figure 10) [2,3]. The CaSO4 crystals were mainly deposited along the regularly distributed convex lines formed on the membrane sheet during
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manufacturing process while the reddish-brown deposits were observed on the entire membrane surface. EDS analysis indicated that the surface deposits primarily consisted of an iron-containing portion (colloidal) and CaSO4 (Figure 11). We later determined that this iron deposit was caused inadvertently by the installation of a new section of pipe during the modification of the flow system
80
Study Of Membrane Fouling And Cleaning In Spiral Wound Modules Using Ultrasonic Time-Domain Reflectometry - Greenberg
that originally utilized plastic or stainless steel components. We believe that the introduction of air into this aqueous, open loop system caused oxidation of iron ions (Fe+2 ---> Fe ÷3) and the subsequent formation of ferric hydroxide. Previous work has shown that insoluble ferric hydroxides (colloidal) can produce fouling of RO systems [12]. The presence of this mixed foulant provided an unanticipated degree of complexity to the experiments and also accounts for the unexpected shift factor response previously described. Formation and deposition of the iron-containing foulant on the membrane surface undoubtedly occurred during the initial exposure of the new section of pipe to water (cycle 1). Fouling due to the deposition of iron-containing layers most likely became more extensive during phase B because the introduction of CaSO4 would facilitate electrolytic corrosion of the iron pipe [ 13]. Correspondingly, QA evidenced only a slight increase during the cycle 1 water-equilibration phase but a much more pronounced increase during the subsequent fouling phase (Figure 7). In contrast, QT was relatively unchanged throughout cycle 1. This behavior would be expected from the deposition of a very thin layer with a relatively loose structure since QA is sensitive to the topographical characteristics whereas QT is most responsive to the thickness of a sufficiently thick deposit. In the present case this suggests that QA and QT respond preferentially to the deposition of the ironcontaining and CaSO4 layers, respectively. If correct, this indicates that the deposition of the iron-containing colloidal material accounted for a significant portion of the overall fouling that occurred during phase B of cycle. This hypothesis is supported by the results of the retentate concentration measurements that indicated values of only 1.6 g/L at the end of phase B. As previously described, the permeation rate and QA responses during the cleaning phase of cycle 1 were puzzling since the former returned to its reference value but the latter decreased only slightly. However, such behavior would be consistent with the fact that while the pure water quickly cleaned the CaSO4, the cleaning protocol was not effective in removing the iron-containing layer. These results also suggest that the iron-containing layer had only a relatively small effect on the permeate flux due to its non-dense colloidal structure. During cycle 2, QA remained relatively stable during phase A in contrast to the increase observed during cycle 1 (Figure 7). The behavior of QA in conjunction with the relatively constant values of QT during the waterequilibration phase of both cycles suggests that any additional iron-based foulinglayer deposition during phase A was not extensive. Results obtained during the cycle 2 fouling phase, when a higher concentration of CaSO4 (1.6 g/l) was used in the feed solution, are consistent with a fouling layer that consisted of more significant amounts crystalline CaSO4 as well as the iron-containing colloidal material. In this case the permeation rate decreased and both shift factors increased. During the subsequent pure-water cleaning phase, the permeation rate and QT return to their respective reference values. 81
Study Of Membrane Fouling And Cleaning In Spiral Wound Modules Using Ultrasonic Time-Domain Reflectometry - Greenberg
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Figure 10. Representative micrographs of the fouling deposits on the membrane surface of module I. Left: Low magnification view showing the calcium-sulfate regions (two dark gray ovals along the picture diagonal) and the iron-containing regions (surrounding widely distributed light areas); Right: Higher magnification view of the needle-like structures characteristic of crystalline calcium sulfate.
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Figure 11 . Representative results obtained from energy-dispersive spectroscopy of the fouling deposits from module I. The backscattered image indicates the specific area from which the iron, calcium and sulfur scans were obtained. The relative brightness of the image indicates the concentration of the respective element. The images indicate that the composition of the oval areas is consistent with calcium sulfate whereas the deposits in the surrounding regions contain significant iron.
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Study Of Membrane Fouling And Cleaning In Spiral Wound Modules Using Ultrasonic Time-Domain Refleetometry - Gteenberg
However, although QA clearly decreases, the reference value is not attained by the end of the cleaning phase. These responses are consistent with the removal of the CaSO4 scalant but the continued presence of the ironcontaining layer. An interesting aspect of the phase C shift factor responses is that both QA and QT evidence a rather sharp decrease after approximately 20 hr of pure water cleaning. This behavior is in contrast to that of the permeation rate, which evidences a more gradual increase over this same time period. These characteristics underscore a fundamental difference between the two types of measurements: whereas the permeation rate represents the average condition of the whole module, the acoustic responses are site specific, i.e., they indicate the condition of a relatively small portion of the module. For example, the focused transducers used in this study have a "spot" size of only a few millimeters in the focal plane. Similarly, the failure of an acoustic parameter to return to the appropriate reference value is consistent with the continued presence of a fouling layer at that specific location even though the permeation rate indicates that the reference condition has been achieved. Moreover, differences in the shift factor responses obtained from transducers attached at different locations on the module can indicate that conditions at these locations are not the same. In the present study, gravimetric measurements indicated that the iron-containing deposit was concentrated on the upstream end of the module, whereas the crystallized CaSO4 scalant was concentrated on the downstream end. This situation accounts for the differences in the responses from the midstream and downstream transducers whereby the former reflects the greater influence of the iron-containing deposit, while the latter is more response to the CaSO4 scalant. The ability to obtain information from multiple locations represents an important advantage of the acoustic measurements since they can provide additional insight regarding module performance. In order to evaluate the reproducibility of the results, a second module was tested using the protocol described in Table 1. Although similar results were obtained for the first cycle in which the lowest concentration of CaSO4 (1.2 g/L) was used, the acoustic signals obtained during the second and third cycles were much noisier than those from module I. We believe that this problem was due to inconsistencies in remounting the transducers that in tuna produced excessive vibrations during module operation. Despite these difficulties, the basic trends described for module I were still evident for module II. As indicated in Table 1, the experimental protocol for module II was modified somewhat in that the experiment was ended and the module was opened after the completion of the cycle 3 cleaning phase. As shown in Figure 12, microscopic examination confirmed the presence of a reddish-brown
85
Study Of Membrane Fouling And Cleaning In Spiral Wound Modules Using Ultrasonic Time-Domain Reflectometry - Greenberg
deposit on the membrane surface; however, here there was no evidence of the white crystalline deposit usually indicative of CaSO4 fouling. The presence of this iron-containing deposit on the surface after cleaning correlates well with the behavior of the shift factors during this cycle of module II operation whereby the factors decreased from their "fouled" values but did not reach their "clean" (reference) values. In contrast, the permeation rate did return to its reference value after the cleaning cycle. Previous studies have shown that at a pH value close to the isoelectric point of ferric hydroxide, fouling occurred via the deposition of loosely packed large aggregates that did not result in a significant decrease in membrane flux [11]. The micrographs shown in Figure 12 indicate the ferric deposits are loosely packed on the membrane surface. Consequently, the permeation rate completely recovers, whereas as the shift factors do not. In this situation the ultrasonic technique appears more sensitive to fouling layer formation and removal than the fluxdecline behavior. Based upon the previous successful UTDR studies of flat sheet membranes, the selection of calcium sulfate as a model fouling system for a first UTDR investigation of fouling and cleaning in spiral-wound membrane modules is a reasonable choice. However, calcium sulfate scaling is generally not considered to be the most common problem in commercial separations. Indeed, fouling due to the deposition of a colloidal layer is a much more common phenomenon, and can cause flux decline in almost all membranebased liquid separations [2]. The rate at which such flux decline occurs will depend upon the specific system and operating conditions and the particular characteristics of the resulting fouling layer. Within this framework, the capability of the UTDR technique to detect mixed-mode fouling comprised of colloidal as well as inorganic deposits is of considerable significance with respect to the application of this technique in many commercial situations in which mixed fouling occurs. 5. CONCLUSION This research continues the efforts to extend the UTDR technique to the measurement of fouling and cleaning in commercial membrane modules. The results of this study indicate that the new signal-analysis methodology along with corresponding software for real-time monitoring is an effective approach for quantitative detection of fouling and cleaning in spiral-wound modules. This research provides the first evidence that the UTDR technique is capable of detecting mixed-mode fouling comprising colloidal as well as inorganic deposits. Overall, the insights obtained from this study provide a strong basis for the continued development of the UTDR technique with an ultimate objective of successfully monitoring membrane fouling and cleaning under the 86
Study Of Membrane Fouling And Cleaning In Spiral Wound Modules Using Ultrasonic Time-Domain Reflectometry - Greenberg
!bt
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Figure 12. Representative micrographs of the membrane surface after fouling and water cleaning: (a) section from module I indicates the presence of both calcium-sulfate crystals and grainy reddish-brown iron deposits after the cycle 3 fouling phase; (b) section from module II shows that only iron deposits remain after cycle 3 cleaning phase; (c) another section from module I shows the presence of both deposits after the cycle 3 fouling phase; and (d) same section as (c) but after overnight immersion in w a t e r - the calcium-sulfate deposits are removed but the iron deposits remain on the membrane surface.
rigors of commercial operating conditions. The research emphasizes the complexity of the fouling process and the necessity of a robust experimental design. In retrospect this study would have benefited from a larger scale effort involving the simultaneous operation of multiple modules, each outfitted with appropriate acoustic hardware. Although difficult to accomplish within cost and space limitations, such an approach represents a more efficient way to successfully evaluate all of the factors that influence the ultrasonic measurements during fouling and cleaning. Despite these limitations, the results clearly indicate the potential advantages of UTDR measurement including enhanced sensitivity and
87
Study Of Membrane Fouling And Cleaning In Spiral Wound Modules Using Ultrasonic Time-Domain Reflectometry - Greenberg
additional insight regarding module conditions as compared with information obtained from permeate measurements alone. ACKNOWLEDGEMENT
The authors gratefully acknowledge the support of this research by the NSF Membrane Applied Science and Technology (MAST) Center at the University of Colorado. REFERENCES [1] [21
[3] [4] [5] [6] [7] [8] [9] [10] [11] [12] [13]
G. Belfort, in: G. Belfort (Ed.), Synthetic Membrane Processes, Chapter 1, Academic Press, Orlando, 1984. A. P. Mairal, A.R. Greenberg, W.B. Krantz, L.J. Bond' J. Membr. Sci. 159 (1999) 185-196. A. P. Mairal, A.R. Greenberg, W.B. Krantz, Desalination 130 (2000) 45-60. L.J. Bond, G.Y. Chai, A.R. Greenberg, W.B. Krantz, Method and Apparatus for Determining the State of Fouling/Cleaning of Membrane Modules, U.S. Patent 6,161,435, December 19, 2000. E. Fountonkidis, Z.B. Maroulis, D. Marinos-Kouris, Desalination 79 (1990) 47-63. M. Brusilovky, J. Borden, D. Hasson, Desalination 86 (1992) 187-222. A. Malik, F.A. Aleem, Desalination 96 (1994) 409-419. F.H. Butt, F. Rahman, U Baduruthanal, Desalination 109 (1997) 323-336. G. Y. Chai, A.R. Greenberg, W.B. Krantz, Euromembrane 99 2 (1999) 357-359. G. Y. Chai, W.B. Krantz, A.R. Greenberg, "Measurement of Membrane Fouling in Spiral-Wound Modules," 6th World Congress of Chemical Engineering, Melbourne, Australia, September 2001. A. P. Mairal, Ph.D. dissertation, University of Colorado at Boulder, 1998. K. Scott, Handbook of Industrial Membranes, 1st ed., Elsevier Advanced Technology, Oxford, UK, 1995. P.L. Mangonon, The Principles of Materials Selection for Engineering Design, Prentice Hall, Upper Saddle River, NJ, 1999.
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New Insights into Membrane Science and Technology: Polymeric and Biofunctional Membranes D. Bhattacharyya and D.A. Butterfield (Editors) © 2003 Elsevier Science B.V. All rights reserved.
Chapter 5
Nonintrusive characterization of fluid transport phenomena in hollow-fiber membrane modules using MRI: An innovative experimental approach C. K. Poh 1, P. A. Hardy 2, Z. Liao 1, W. R. Clark 3, and D. G a o 1'2'4. 1Department of Mechanical Engineering, University of Kentucky, Lexington, KY, USA. 2Center for Biomedical Engineering, University of Kentucky, Lexington, KY, USA. 3Renal Division, Baxter Healthcare Corporation, McGaw Park, IL, USA. 4Department of Thermal Science and Energy Engineering, University of Science and Technology of China, Hefei, Anhui, China *Corresponding author, Phone: (859) 257-6336 ext. 80666, Fax: (859) 2573304, e-mail:
[email protected] ABSTRACT A nonintrusive experimental tool is useful for a better physical understanding of fluid transport phenomena in hollow-fiber membrane modules and as test cases for validating, improving, or developing numerical models. In this chapter, we introduced two innovative, nonintrusive flow-imaging techniques using magnetic resonance imaging (MRI) for characterizing fluid transport phenomena in hollow-fiber membrane modules. These flow-imaging techniques are called the 2-D Phase-Contrast (2DPC) and 2-D FourierTransform (2DFT) techniques. The principles, validation, advantages, limitations, and some examples of experimental results are presented. We used the 2DPC technique to study the spatial flow distribution and the 2DFT technique to characterize the flow profile and quantify the local ultrafiltration rates in hollow-fiber artificial kidneys (also known as hemodialyzers). These flow-imaging techniques are equally applicable to other hollow-fiber membrane modules. 1.
INTRODUCTION
Knowledge of experimental flow distribution in hollow-fiber membrane modules provides a valuable insight into the real fluid transport phenomena in these modules, which cannot be predicted by numerical models, and it may lead
89
Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
to an improved engineering design of these modules. Any flow maldistribution such as shunt or occluded flow will reduce the separation performance of these modules. Ideally, a uniform flow distribution is preferred in both the shell and lumen sides of these modules for optimal separation performance. Nonintrusive flow visualization in hollow-fiber membrane modules is impossible by the use of conventional flow visualizing tools. Conventional flow visualizing tools are intrusive, where they require seeding particles or probes to be inserted into the flow, and have limited spatial resolution. The small inner diameter of the hollow fibers (200 to 500 ~tm in hemodialyzers, 100 to 400 ~tm in permeators, and 200 to 400 ~tm in contactors) and the presence of large number of hollow fibers (10,000 to 12,000 in hemodialyzers, 300 to 3,500 in permeators, and 900 to 12,500 in contactors) in a given module complicates the flow distribution and poses a great challenge in visualizing the flow in these modules. In most cases, small-scale countercurrent flow is realized in these modules due to the flow inside the hollow fibers (lumen-side flow) moves in the opposite direction to that outside the hollow fibers (shell-side flow). Conventional flow visualizing tools, with limited spatial resolution, are incapable of capturing this small-scale countercurrent flow. Recent advancement in MRI-based flow imaging technique has made it an attractive, nonintrusive tool for studying the fluid transport phenomena in hollow-fiber membrane modules. However, the current application of MRI in the study of fluid transport phenomena in hollow-fiber membrane modules remains minimal. Hammer et al. [1], Heath et al. [2], and Donoghue et al. [3] used MRI to investigate flow distribution in hollow-fiber bioreactors. Pangrle et al. [4], Zhang et al. [5], Laukemper-Ostendorf et al. [5], Osuga et al. [7], and Poh [8] used MRI to study flow distribution in hollow-fiber hemodialyzers. In this chapter, we introduced two innovative, nonintrusive flow-imaging techniques called the 2-D Phase-Contrast (2DPC) and 2-D Fourier-Transform (2DFT) techniques. We used the 2DPC technique to investigate the spatial flow distribution and the 2DFT technique to characterize the flow profile and quantify the local ultrafiltration rates in hollow-fiber hemodialyzers [8-14]. Some examples of these experimental results are presented in this chapter to illustrate the capabilities of these flow-imaging techniques. An extensive discussion on the principles of MRI is available in a number of books [ 15-18]; therefore, it will not be presented in this chapter. We will only discuss the commonly used pulse sequences in MRI. In addition, we will discuss the principles, advantages, limitations, validation, and applications of the 2DPC and 2DFT techniques in hollow-fiber membrane modules. Other MRI-based flow imaging techniques are also available, such as the Delays Alternating with Nutations for Tailored Excitation (DANTE) [19], Multiple Overlapping ThinSlab Acquisition (MOTSA) [20], Real-Time Acquisition and Velocity Evaluation (RACE) [21], and n Echo-Planar Imaging (PEPI) [22]. The 90
Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
discussion of these flow-imaging techniques is beyond the scope of this chapter. In general, different pulse sequences of MRI can be designed or modified for different flow-imaging techniques. 2.
ADVANTAGES OF THE 2DPC AND 2DFT TECHNIQUES
The 2DPC and 2DFT techniques introduced in this chapter are nonintrusive or nondestructive, where they do not require any seeding particles or insertion of probes. Therefore, these flow-imaging techniques do not disturb the natural occurrence of the flow in hollow-fiber membrane modules. The nonintrusive nature of these flow-imaging techniques allow us to visualize and quantify the natural spatial and temporal behavior of the flow in hollow-fiber membrane modules that is otherwise impossible by means of conventional experimental tools such as hot-wire sensors. These flow-imaging techniques also do not require the object of interest to be transparent; hence, they are particularly advantageous for investigating fluid transport phenomena in hollowfiber membrane modules because these modules are usually opaque, where optical fluid visualizing tools cannot be used. Experimental data acquired from these flow-imaging techniques can be used as test cases to develop, validate, or improve realistic numerical models for use in predicting the separation performance of hollow-fiber membrane modules under different engineering designs and operating conditions. Countercurrent flow is commonly used in hollow-fiber membrane modules to optimize their separation performance. The 2DFT technique is capable of measuring simultaneously countercurrent flow in these modules that is otherwise impossible by means of other nonintrusive fluid-visualizing tools such as the Laser Doppler Velocimetry (LDV) and Particle Imaging Velocimetry (PIV). The LDV and PIV are incapable of measuring countercurrent flow in hollow-fiber membrane modules due to the presence of several thousands of hollow fibers with a relatively small inner diameter of about 100 to 500 ~tm. These hollow fibers allow countercurrent flow to occur simultaneously in a small region of about 250 by 250 btm2, where one fluid flows inside the lumen and the other fluid flows countercurrently outside the lumen of these hollow fibers. The LDV and PIV are incapable of even measuring unidirectional flow in such a small flow region or small-scale flow field [23]; therefore, they are incapable of measuring countercurrent flow in such a small-scale flow field. As in all areas of fluid mechanics, the nonintrusive 2DFT technique, capable of measuring countercurrent flow in such a small-scale flow field, is of great value for studying the flow distribution in hollow-fiber membrane modules.
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Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
3.
OVERVIEW OF STANDARD IMAGING PULSE SEQUENCES
The magnetic resonance signal emitted from a fluid of interest is a function of several imaging parameters such as the proton density, T1- and T2relaxation times, chemical shift, and fluid motion. The Tl-relaxation time is the time for the longitudinal magnetization to grow from zero to 63 % of its initial value, and the T2-relaxation time is the time for the transverse magnetization to decay to 37 % of its initial value. The relative contributions of each imaging parameter determine the final quality of the reconstructed MR images, and they are controlled by the radio frequency pulses, magnetic-field gradient pulses, and data acquisition timing. A pulse sequence is a set of radio frequency and magnetic-field gradient pulses used in conjunction with the data acquisition timing to produce the desired MR images. The two most commonly used imaging pulse sequences in the clinical MRI are the spin-echo and gradient-echo pulse sequences. The time for a complete cycle of a pulse sequence is called the repetition time, TR, and the time between the first maximum echo signal emitted is called the echo time, TE.
3.1.
Spin-Echo Pulse Sequence The standard spin-echo pulse is shown in Figure 1. The radio frequency pulses are denoted as RF, the slice-selective gradient pulses are denoted as Gstice,the phase-encoding gradient pulses are denoted as Gehase,and the frequency-encoding gradient pulses are denoted a s GReadin Figure 1. The magnetic resonance signal, an echo signal that emitted from the fluid of interest after the application of the 90-degree and 180-degree radio frequency pulses is shown on the last line in Figure 1. The frequency-encoding gradient is denoted a s aReadbecause it is frequently referred to as the read-out gradient. The spatial information in the MR images comes from the frequencyencoding gradient and phase-encoding gradient. By convention, the main static magnetic field in a MR imager is always aligned with z-axis. The frequencyencoding gradient is applied along the x-axis and is commonly denoted as Gx instead of GReta. The phase-encoding frequency is applied along the y-axis and is commonly denoted as Gy instead of Gpha~e. The frequency-encoding gradient creates a position-specific frequency distribution along the x-axis, and the phase-encoding gradient creates a positionspecific phase distribution along the y-axis. To create a MR image with Nx by Ny pixels, we need to sample Nx data points from the frequency-encoding gradient and to change the phase-encoding gradient Ny times with different gradient values. A MR image is then reconstructed from the recorded raw data, containing both position-specific frequency and position-specific phase distribution, through an inverse two-dimensional Fourier transformation where 92
Nonintrusive Characterization O f Fluid Transport P h e n o m e n a In Hollow-Fiber M e m b r a n e Modules Using MRI: A n Innovative Experimental A p p r o a c h - Poh
90"
180"
90' .................
I
GSlice I
I
GPhase
GRead 0 I
Signal
.................
½TE v
,_.. ~-~
I
'
TR
I
.--II
v I
I
DataSampling Figure 1. The standard spin-echo pulse sequence. Reprinted with permission from Poh [8]. the acquired signal is assigned to the appropriate pixel in the MR image. The acquired raw data set is commonly called the k-space. In a standard spin-echo pulse sequence as shown in Figure 1, the sliceselective gradient, together with the 90-degree radio frequency pulse, is used to select an image slice. The image slice thickness is controlled by the magnitude of the slice-selective gradient, and the bandwidth and shape of the radio frequency pulse. The steeper the slice-selective gradient or the smaller the bandwidth of the radio frequency pulse, the smaller the image slice thickness. The 90-degree radio frequency pulse shown in Figure 1 is also used to flip the net magnetization of nuclear spins into the transverse plane. Prior to the 180degree radio frequency pulse, a stepped phase-encoding gradient is applied to encode the spatial information along the y-axis. After the 90-degree radio frequency pulse, the transverse magnetization of nuclear spins started to decay
93
Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
due to the observed transverse (T2*)relaxation. T2* relaxation is the dephasing of nuclear spins caused by the spin-spin interactions and magnetic field inhomogeneity. A 180-degree radio frequency pulse is then applied at exactly one-half TE. The 180-degree radio frequency pulse reverses the dephasing effect of nuclear spins, caused by the magnetic field inhomogeneity, by flipping them in a mirror-like fashion. The 90-degree and 180-degree radio frequency pulses are sometimes called the excitation and refocusing pulses, respectively. After the 180-degree radio frequency pulse, the rephased nuclear spins cause the transverse magnetization to grow to a maximum value, and then it decays again due to the T2 relaxation instead of T2* relaxation. The dephasing effect of nuclear spins due to magnetic field inhomogeneity is time independent, and that due to spin-spin interactions is time dependent. The 180-degree radio frequency pulse corrects the dephasing effect due only to magnetic field inhomogeneity, but not that due to spin-spin interaction. Therefore, after the 180-degree radio frequency pulse, the recovered transverse magnetization experiences T2 relaxation instead of T2* relaxation because the dephasing effect due to magnetic field inhomogeneity has been corrected by the 180-degree radio frequency pulse. The recovery of the transverse magnetization by a 180-degree radio frequency pulse is called a spin echo, and it generates an echo signal as shown in Figure 1. This echo signal is acquired in the presence of the frequencyencoding gradient. The frequency-encoding gradient provides spatial information along the x-axis. A similar frequency-encoding gradient is applied before the 180-degree radio frequency pulse to ensure that any stationary nuclear spins present are rephased during the spin echo. This process is repeated Ny times using different values of the phase-encoding gradient to acquire Ny data points along the y-axis. The frequency-encoding gradient lobe applied after the 180-degree radio frequency pulse has twice the area of the frequency-encoding gradient lobe applied before the 180-degree radio frequency pulse, allowing the full echo to be refocused. The advantage of the spin-echo pulse sequence is that it is independent of magnetic field inhomogeneity due to the use of a 180-degree refocusing radio frequency pulse. On the down side, however, the time TR in this pulse sequence is relatively long, typically 200 to 2000 ms. Therefore, the spin-echo pulse sequence generally requires a long image-acquisition time. Its TR is relatively long due to the use of a 180-degree radio frequency pulse and to the time required for the longitudinal magnetization to recover before the application of another 90-degree radio frequency pulse. The reason for waiting for the longitudinal magnetization to recover before the application of another 90degree radio frequency pulse is to make sure that sufficient amount of signal is emitted when subsequent 90-degree radio frequency pulses are applied.
94
Nonintrusive Characterization Of Fluid TransportPhenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
Alternative to the spin-echo pulse sequence is the gradient-echo pulse sequence that uses a smaller flip angle (< 90°), and this pulse sequence does not require a 180-degree refocusing radio-frequency pulse. 3.1.2. Gradient-Echo Pulse Sequence The standard gradient-echo pulse sequence is shown in Figure 2. In a gradient-echo pulse sequence, a radio frequency pulse with a smaller flip angle (< 90 °) is applied to flip the net magnetization of nuclear spins compared to that applied in the spin-echo pulse sequence. In addition, a bipolar frequencyencoding gradient pulse instead of a 180-degree radio frequency pulse is used to rephase or refocus the dephasing nuclear spins after the first radio-frequency pulse compared to that in the spin-echo pulse sequence. The gradient-echo pulse sequence is also known as the gradient-recalled pulse sequence. Initially, a radio frequency pulse is applied to flip the net magnetization of nuclear spins by an angle, a with respect to the axis of the applied static magnetic field, where a is usually between 0 and 90 °. This radio frequency is commonly denoted as an a pulse. Similar to the spin-echo pulse sequence, the image slice is selected using the slice-selective gradient together with the a pulse as shown in Figure 2. After the application of an a pulse, a stepped phase-encoding gradient is applied. The stepped gradient modulates the phase of the signal along the y-axis, and hence provides the spatial information along the y-axis. At the same time, a bipolar frequency-encoding gradient, as shown in Figure 2, is applied. The first negative lobe of the frequency-encoding gradient dephases the transverse magnetization of the nuclear spins caused by the a pulse. The second positive lobe of the frequency-encoding gradient then rephases the transverse magnetization and hence increases its magnitude. Subsequently, a gradient echo is produced at the time TE. The echo signal is acquired during the rephasing of the transverse magnetization. The frequency-encoding gradient also provides the spatial information along the x-axis. This process is repeated Ny times using different values of the phase-encoding gradient to acquired Ny data points along the y-axis. One of the advantages of the gradient-echo pulse sequence compared to the spin-echo sequence is its short image-acquisition time. With a smaller flip angle, longitudinal magnetization takes shorter time to recover compared to a 90 ° flip angle, and the pulse sequence requires a shorter TE and TR times. Therefore, the gradient-echo pulse sequence has a shorter total imageacquisition time compared to the spin-echo pulse sequence. On the other hand, the gradient-echo pulse sequence is sensitive to magnetic field inhomogeneity. The bipolar frequency-encoding gradient used in the gradient-echo pulse sequence does not reverses the dephasing effect of the magnetic field inhomogeneity as does the 180-degree refocusing radio95
Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
G
RF-Gs,= I GPhase
GR~d
Signal TE TR I
Data Sampling
Figure 2. The standard gradient-echo pulse sequence. Reprinted with permission from Poh [8].
frequency pulse used in the spin-echo pulse sequence. Therefore, the signal generated from the gradient echo in the gradient-echo pulse sequence is weaker than that generated from the spin echo in the spin-echo pulse sequence. 4.
OVERVIEW OF FLOW IMAGING TECHNIQUES
Flow phenomena in MRI can be categorized into the phase shift phenomenon and time-of-flight (TOF) phenomenon [24]. The phase shift phenomenon is used in the 2-D Phase-Contrast (2DPC) technique; a combination of phase shift phenomenon and TOF phenomenon is used in the 2D Fourier-Transform (2DFT) technique. A full discussion on these flow phenomena and imaging techniques can be found in the references cited throughout the rest of this chapter.
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Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
4.1.
2-D Phase-Contrast Technique The 2-D Phase-Contrast (2DPC) technique incorporates flow-encoding gradients into a standard gradient-echo imaging sequence and makes use of flow-induced phase shift to encode velocity information of a constant flow. The precessional frequency, co of a nuclear spin in a static magnetic field is given by the Larmor equation:
(1)
co = yB o
where ~, is the gyromagnetic ratio of hydrogen atoms (42.57 MHz/Tesla), and Bo is the strength of the applied static magnetic field.
Suppose that the nuclear spin is moving in a spatially homogeneous magnetic field, Bo, and when a magnetic field gradient, G is introduced, the precessional frequency of the moving nuclear spin becomes positional dependent, and it is given by: co(,')= coo + 7G.r
(2)
where COo is the precessional frequency of the nuclear spin in Bo with the absence of the magnetic field gradient, G, ?, is the gyromagnetic ratio, and r is the positional vector of the nuclear spin. The change of phase angle or phase shift, ~ with respect to the time of the moving nuclear spin relative to its initial position at r = 0 can be defined by rearranging Eq. (2): d¢
dt
= co(*')-coo =7G'r
(3)
Integrating Eq. (3) with respect to time, t, we have: ¢(t)= Y ~G(t).r(t)dt
(4)
where ~ (t) is the phase shift of the nuclear spin as a function of time relative to its initial position at r = 0. In addition, the positional vector of the nuclear spin, r can be expressed in terms of time derivatives of its value at t = 0 using the Taylor series expansion:
97
Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
"(') = "o +
,=o t + - =L J,=o
= ro + Vot + laot2
+"" + "!l or" Jt
=o
(5)
+ ...
where to, Vo, and ao are the position, velocity, and acceleration vectors of the nuclear spin, respectively, at an initial time, t = 0. Substituting Eq. (5) into Eq. (4), we have: 1
m • j'c(t)a, +,,o. j'c(t),a, + r,o.
+...
(6)
Alternatively, Eq. (6) can be written in terms of moments of the magnetic field gradient as:
~b(t)= y M
1 Z
o . r o + 7"M 1 . v o + -~ y M 2 . a o +
o oil
(7) and
M i - ~G(t)tidt
where M o , M 1 , and M 2 are the zeroth, first, and second moments of the magnetic field gradient, respectively. The higher order terms (+...) in Eq. (7) are usually very small and are negligible. The magnetic-field gradient moments in Eq. (7) can be manipulated accordingly during the experiment by choosing the proper shape of the magnetic-field gradient waveform. For instance, the zeroth moment of the magnetic field gradient can be set to zero by choosing a magnetic field gradient that has two opposite gradient lobes with equal areas as a function of time, called the bipolar gradient [30]. In theory, any moments can be set to zero, and we can generate either velocity images by setting the zeroth moment of the magnetic field gradient to zero or acceleration images by setting the zeroth and first moments of the magnetic field gradient to zero. For a uniform flow i.e. without acceleration, Eq. (8) can be rewritten as: (9)
~(t)-- ~ M 0 "1"o + ~M1 "Vo
911
Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
My G(t)
g Spin
I'
~
Phase
1st GradientPulse
1
Stationary'7 Mx
2nd GradientPulse
¢~sl
9
Position
¢~s2!
StationarySpin
MovingSpin
Figure 3. The effect of a bipolar gradient pulse on the phase shifts of stationary and moving spins. Reprinted with permission from Poh [8].
The first term on the fight-hand side of Eq. (9) is proportional to the position and area under the gradient waveform and is called the positional encoding term. The second term on the right-hand side of Eq. (9) is proportional to the velocity and is called the velocity encoding term. In the 2DPC technique, to measure the axial velocity of the flow, in this case in the z-direction, we choose a bipolar magnetic-field gradient pulse, or sometimes called a bipolar gradient pulse, Gz(t) that has two opposite gradient lobes with equal areas as shown in Figure 3 such that its zeroth moment is zero (positional independent), i.e. the net phase shift is independent of any stationary nuclear spins. As shown in Figure 3, the bipolar gradient pulse has no effect on the phase shifts induced by the stationary nuclear spin because the these phase shifts at the first and second gradient lobes of the bipolar gradient pulse cancel out each other (¢s2-¢s~ =0). Nevertheless, the phase shifts induced by the moving nuclear spin at the first and second gradient lobes of the bipolar gradient pulse do not cancel out each other (¢M2-¢M~ ~ 0), and the net phase shift, ev, is 99
Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Experimental Approach - Poh
a
RF
Using
MRI:
An
Innovative
a
i
~q~
~~(?:~.....
]") FlowCompensated
i
. . . . ~1 ............
,,
GS,~,tF~
. . . . . . . . . . . . . .
RF
FlowEncoded __1 ............
.............
i i i
rE i
Figure 4. The 2-D Phase-Contrast (2DPC) pulse sequence. The radio frequency pulses are denoted as RF, and the flow-encoding gradient pulses are denoted as Gsl~ce/Flow. The phase-encoding gradient pulses Geh~se and the frequency-encoding gradient pulses Ggead are not shown in this figure. An echo signal (not shown) is generated at the time TE and a is usually between 0 and 90 °. Reprinted with permission from Poh [8].
proportional to the velocity of the moving nuclear spin, where #v =¢M2--~MI" Therefore, by using a bipolar gradient pulse that has two opposite gradient lobes with equal areas, the moving nuclear spin has a net phase shift proportional to its velocity, and this net phase shift is independent of any stationary spins that may have been present in the system. The net phase shift due to the moving nuclear spin in the z-direction, ~,.=(t) is: ~v,z (t) -- y M 1v z + (some error source terms)
100
(10)
Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
The phase shift, #v.z(t) is now velocity dependent and positional independent. This phase shift is sometimes called the flow-induced phase shift. However, this flow-induced phase shift is associated with some error sources or background phase variations such as magnetic field inhomogeneity, imperfections in pulse sequence tuning, and eddy-current effect. The eddy-current effect will be discussed in the latter part of this chapter. Two data sets with different flow sensitivities in an interleaved fashion are acquired during a single experiment to eliminate the background phase variations. The imaging pulse sequence for the 2DPC technique is shown in Figure 4. The first data set acquired is flow compensated [25-28], i.e. flow desensitized, where the first moment of the magnetic-field gradient pulse is set to zero; and the second data set is flow encoded, i.e. flow sensitized, to record the axial velocity (z-direction) of the flow. Flow compensation in the first data set is also known moment nulling. The flow sensitivity in the second data set is controlled by the strength or amplitude of the magnetic-field gradient pulse. Since both data sets have an equal background phase variations for a given experiment, the first data set is used to correct for the background phase variations in the second data set by subtracting the first data set from the second data set. Therefore, the net phase shift, in radian, is given by:
(11)
A ~v,z -- y'AI~l Vz
where the net phase shift, aq~v.,is proportional to the velocity in the z-direction and is independent of background phase variations. AM1 is the difference in the first moment of the z-direction gradient pulses. An important parameter in the 2DPC technique is the encoding velocity, Ve,c. The encoding velocity defines the dynamic range of velocities measurable by the imaging technique. It is controlled by the difference in the first moment of the gradient pulses, zJM1 and is defined as:
Venc
7,~1
(12)
By combining Eq. (11) and Eq. (12), and rearranging; we have the measurable velocity, Vz given by :
(13)
vz=A~v,z venc
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Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
The net phase shift, a#v,z can only take values between -zr to +zr for countercurrent flows, and as a result, the measurable velocity, Vz is between - Venc to + Ve~c. The encoding velocity, Ve#c has to be determined a priori, so that the measurable velocity, Vz is within the velocity range of the flow of interest. Phase shift aliasing will occur when the measured net phase shift exceeds the limits of -n: to +zr (i.e. the measured velocity exceeds the range of the encoding velocity) resulting in a potential ambiguity in determining the flow velocity. The measure net phase shift is then converted into axial velocity using Eq. (13) with a predetermined encoding velocity. The precision of the velocity determined from the measured net phase shift depends on the noise in the phase shift measurement. The standard deviation of the velocity determined from the measured net phase shift, trv is given by:
Ve c
(14)
SNR
where S N R is the signal-to-noise ratio in the acquired image. Eq. (14) shows that the noise or standard deviation of the velocity can be reduced by increasing the S N R or by decreasing the Venc. 4.1.2. 2-D Fourier-Transform Technique The 2-D Fourier-Transform (2DFT) technique, on the other hand, incorporates flow-encoding gradients into a standard spin-echo imaging sequence. This imaging technique makes use of the combined effects of time-offlight (TOF) and phase-shift imaging techniques, proposed by Moran [29], called the flow-velocity zeugmatographic imaging. In addition to the incorporation of flow-velocity zeugmatographic imaging technique into a standard spin-echo pulse sequence as introduced in this chapter, the flowvelocity zeugmatographic imaging technique can be incorporated into any MR imaging pulse sequences without interfering with the spatial encoding and reconstruction algorithm of the imaging technique [29]. Basic understanding of the TOF phenomenon is required for the 2DFT technique. Reviews on TOF phenomenon can be found in [24,30-34]. In a standard spin-echo pulse sequence, TOF phenomenon demonstrates two flow effects: the inflow and outflow effects. In this 2DFT technique, we make use of the inflow effect from the TOF phenomenon. Therefore, only the inflow effect will be discussed in this chapter. Inflow effect refers to a moving of fresh or unexcited nuclear spins through an imaging slice perpendicular to the flow direction. Since these fresh 102
Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
nuclear spins have not experienced any radio frequency pulses before, they emit a stronger signal when they are first excited by a radio frequency pulse in the imaging slice than any surrounding saturated stationary nuclear spins. The stationary nuclear spins within the imaging slice have experienced several excitation pulses, and the TR time is too short for their longitudinal magnetization to recover. Therefore, these stationary nuclear spins are either partially or fully saturated, and they emit either very little or no signal at all during the appearance of the spin echo. The inflow effect is dependent on the timing of the radio frequency pulse (TR and TE) and the velocity of the flow. In a given imaging slice, for moving nuclear spins to emit a maximum signal during the spin echo or data acquisition, these moving nuclear spins must experience both the 90-degree and 180-degree pulses in a standard spin-echo pulse sequence and move perpendicular to the imaging slice. Any nuclear spins that move through the imaging slice in the time less ½ TE will not be excited by the 180-degree pulse and hence will not emit any signals. The maximum velocity allowable, Vmaxfor a moving nuclear spin to experience both the 90-degree and 180-degree pulses for a standard spin-echo pulse sequence is given by: V~a~ = Z~Slice TR
(15)
where Zst~ceis the thickness of the image slice, and TR is the repetition time. In general, the signal emitted from the moving nuclear spins increases linearly with increasing velocity [24], provided the velocity is less than Vmax.For nuclear spins that move faster than Vm~x,a significant fraction of the moving nuclear spins may have experienced the 180-degree pulse before the 90-degree pulse. Since these fast moving nuclear spins have no transverse magnetization, they do not contribute to the spin echo, and the signal will eventually decreases as velocity increases beyond Vmax. A detailed analysis by Wehrli [24] shows that four different populations of moving nuclear spins can coexist in an imaging slice of a standard spin-echo pulse sequence. The moving nuclear-spin populations are fresh nuclear spins, fully relaxed nuclear spins, inverted nuclear spins, and saturated nuclear spins as shown in Figure 5. Assuming a plug flow with velocity, v in a circular pipe and taking an imaging slice with slice thickness, Zst~ce as shown in Figure 5. A 90degree pulse is applied at the time, t = 0, and a bolus of fluid is tracked with increase in time. At the time t = ½ TE, the 180-degree pulse is applied, followed by another 90-degree pulse at the time, t = TR. Finally, at the t = TR + ½ TE, that is at the time the 180-degree pulse is applied after the application of the second 90-degree pulse, the bolus of fluid has traveled a total distance of v(TR + ½ TE), 103
Nonintrusiv¢ Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using M R I : Experimental Approach - Poh
II t t !
I I I
!
t
An Innovative
t
!i ................... , v(rm) t
i
t
v(m)
i
v(TR * 'A TE}
Figure 5. Four different populations of nuclear spins in an imaging slice of a plug flow after a 90°-180°-90°-180 * spin-echo pulse sequence cycle. Region 1 represents fresh nuclear spins, region 2 represents fully relaxed nuclear spins, region 3 represents inverted nuclear spins, and region 4 represents saturated nuclear spins. Zslice is the imaging slice thickness, TR is the repetition time, TE is the echo time, and v is the flow velocity. Reprinted with permission from Poh [8].
and fresh nuclear spins have entered into the imaging slice. In addition, the signal intensities from the moving nuclear-spin populations that emit signal during the spin echo can be determined from the Bloch equations [24]: S Relaxed = Mt=oe (-TE//T2)
Slnverted = Mt= 0 1 - 2e k
T1
.I e(-TE//T2)
(16)
Ssaturate d = M t = 0 1 -
2e L
--
+e(_rR//n) e(-rE//~2)
104
Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
where SRelaxed, Slnverted, and Ssaturated a r e the signal intensities for the fully relaxed, inverted, and saturated nuclear spins respectively, M ~ is the magnetization of nuclear spins at time, t = 0, TE is the echo time, and TR is the repetition time. In the 2DFT technique, the velocity information is encoded by a stepped flow-encoding gradient as shown in Figure 6. A 90-degree radio frequency pulse is initially applied followed by a 180-degree radio frequency pulse at ½ TE. An echo signal is generated after the 180-degree radio frequency pulse. TE is the time between the first 90-degree radio frequency pulse and the maximum echo signal. TR is the time between successive 90-degree radio frequency pulses. ~ is the period of the stepped flow-encoding gradient pulses. During data acquisition, unlike in the 2DPC technique where the flowencoding gradient pulse is constant, the flow-encoding gradient in the 2DFT technique is stepped N times, and the resulting velocity data is Fourier transformed into N velocity images, each with a specific velocity bin size, Av. As the name of the imaging technique implied, these N velocity images are the Fourier-velocity images of the flow, and each Fourier-velocity image covers a specific velocity range or bin size. The velocity bin size, Av of the Fouriervelocity images reconstructed from the 2DFT technique is controlled by the thickness of the image slice, number of reconstructed Fourier-velocity images, and period of the stepped flow-encoding gradient pulses as shown by the following equation:
A v = Z Slice
(17)
Nr
where Zst~ is the thickness of the image slice, N is the number of reconstructed Fourier-velocity images, and r is the period of the stepped flow-encoding gradient pulses. In contrast to the 2DPC technique, which encodes velocity in the phase images, the 2DFT technique encodes velocity in the reconstructed Fourier-velocity images as magnitude images. The 2DFT technique is less sensitive to phase dispersion within a single pixel compared to the 2DPC technique and hence is more accurate in measuring less steady flow. Phase dispersion within a single pixel is due to the presence of a velocity distribution in a single pixel. The 2DPC technique, on the other hand, is extremely sensitive to phase dispersion within a single pixel because this flow imaging technique makes use of flow-induced phase shift to encode velocity information. On the contrary, the 2DFT technique requires a longer total imageacquisition time compared to the 2DPC technique due to the requirement of multiple flow-encoding steps in the 2DFT technique to reconstruct multiple Fourier-velocity images. 105
Nonintrusive Characterization Of Fluid Transport Phenomena E x p e r i m e n t a l A p p r o a c h - Poh
90" |
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e
VALIDATION OF THE 2DPC AND 2DFT TECHNIQUES
We validated the accuracy of the 2DPC and 2DFT techniques using a flow phantom or test flow in a Siemens 1.5T MAGNETOM Vision whole-body MR imager (Siemens AG, Erlangen, Germany). The MR imager is identical to that used for human imaging. It has a maximum gradient strength of 25 mT/m and a rise time of 600 ~ts. All image processing was performed using userwritten MATLAB programs together with the MATLAB Image Processing Toolbox 2.2 software (The MathWorks Inc., Natick, MA, USA) rtmning on a Linux workstation. 106
Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
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Figure 7. The experimental setup of a flow phantom. Reprinted with permission from Poh [8].
5.1.
Experimental Methods The flow phantom was made of countercurrent laminar flows in two identical Plexiglas tubes at a steady flow rate as shown in Figure 7. The Plexiglas tubes were 1.8 m in length with inner and outer diameters of 1.6 and 1.9 cm, respectively. These Plexiglas tubes were positioned parallel to each other and placed along the axis of the main magnetic field inside the bore of the MR imager. A solution made of distilled water doped with cupric sulfate pentahydrate, CuSO4"5H20 (ACS Reagent, Sigma Chemical Co., St. Louis, MO, USA), with a final concentration of three millimolar (mM), was gravimetrically fed to one of the tubes from an elevated constant-head reservoir. A calibrated rotameter (Cole-Parmer Instrument Co., Vernon Hills, IL, USA) was used to control the desired flow rate. The flow out of the tube was channeled back to the other tube to create a countercurrent flow. The purpose of adding cupric sulfate pentahydrate to the distilled water was to increase the signal-to-noise ratio and to reduce the total image-acquisition time of the experiment. A peristaltic pump (Cardiovascular Equipment, Sams Inc., Ann Arbor, Michigan, USA) was then used to recycle the flow out of the flow phantom into the elevated constant-head reservoir. The flow rate used in validating the accuracy of the flow-imaging techniques was 100 mL/min as measured from the calibrated rotameter. This flow rate corresponds to a Reynolds number of 135 and a mean velocity of 0.9 cm/s, assuming a laminar pipe flow, in each Plexiglas tube.
107
Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
All imaging was performed, at the mid-length cross sections, perpendicular to the flow direction, of the Plexiglas tubes (Figure 7). Long Plexiglas tubes (1.8 m) were used to ensure that fully developed laminar, countercurrent flows occur at the mid-length of these tubes. In addition, two test tubes filled with the similar prepared solution were placed beside the Plexiglas tubes at the imaging plane. These test tubes served as a reference for stationary flow. All velocities were measured in the axial direction; the direction along the length of the Plexiglas tubes. The flow phantom was first imaged using the 2DPC technique with the maximum velocity encoded at 10 cm/s followed by the 2DFT technique. Other imaging parameters for the 2DPC technique were TR = 45 ms, TE = 12 ms, a = 15 °, and slice thickness = 4 mm. TR is the repetition time, TE is the echo time, and a is the flip angle. The total image-acquisition time was approximately one minute. All phase images acquired had a matrix dimension of 512 x 512 with a field-of-view (FOV) of 300 mm. For the 2DFT technique, twenty-four Fouriervelocity images with velocities ranging from -2.5 to +2.3 cm/s, each with a velocity bin size of 0.2 cm/s, were acquired. Other imaging parameters for the 2DFT technique were TR = 200 ms, TE = 111 ms, r = 100 ms, and slice thickness = 4.8 mm. The total image-acquisition time was approximately ten minutes. All Fourier-velocity images acquired had a matrix dimension of 64 x 256 with a FOV of 50 x 100 mm.
5.1.2. Experimental Results We defined the flow moving into the MR imager as the positive flow and that moving out of the MR images as the negative flow.
Figure 8. A raw phase image of the flow phantom. Reprinted with permission from Poh [8]. 108
Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using Experimental Approach - Poh
Flow Rate = 100 mL/min
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5.2.2. 2DPC Technique Figure 8 shows the raw phase image acquired by the 2DPC technique from the flow phantom. The image artifact was removed during image processing, and the image processing methodology has been discussed in Poh [8]. The processed phase image was then converted into a velocity image using Eq. (13). A 3-D representation of the velocity profile (Figure 9) acquired from the mid-length of the flow phantom shows that the countercurrent flows had the characteristics of a fully developed laminar pipe flow, and the maximum positive and negative velocities were approximately equal in magnitude. These countercurrent flows were fully developed laminar flows because at these cross sections of the tubes, the ratio of the tubes' length to their inner diameters (L/d) was 56. Generally, a laminar pipe flow is fully developed when the L/d ratio [35] is greater than 0.06Re, Re being the Reynolds number determined from the inner diameter of the pipe. In our flow phantom, its 0.06Re had a value of eight,
109
NonintrusiveCharacterizationOf FluidTransportPhenomenaIn Hollow-FiberMembraneModulesUsingMRI:An Innovative ExperimentalApproach- Poh T a b l e 1. Comparison of experimental data measured from the 2DPC technique and rotameter. Reprinted with permission from Poh [8] Positive-Flow Tube in a Flow Phantom
Flow Rate (mL/min) Maximum Velocity a (cm/s)
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% Relative Error b
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2DPC Data - Rotameter Data
x 100 %
Rotameter Data
which was much less than its L/d ratio; therefore, the countercurrent flows at the mid-length of the flow phantom were fully developed laminar flows. In addition to the parabolic velocity profile of the flow phantom, the velocities measured from the 2DPC technique were in good agreement with the analytical velocities determined from the Poiseuille's equation based on the flow rate measured from the calibrated rotameter (Table 1). To further validate the accuracy of the 2DPC technique, we quantified the flow rates and determined the flow areas in each tube from the velocity image. The flow rates in each tube quantified from the 2DPC technique were calculated by multiplying the area of a single pixel with the sum of all velocities in each tube, and they were in good agreement with that measured from the calibrated rotameter (Table 1). The flow areas in each tube, on the other hand, were determined from the velocity image by multiplying the area of a single pixel with the sum of all pixels in each tube; they were 2.04 and 1.96 cm 2 in the tubes with positive and negative flows, respectively, compared to the actual flow area of each tube, 1.96 c m 2. Therefore, the accuracy of the 2DPC technique, with the imaging parameters used in this study, has a six to eight percent relative error in the measured velocity, a three to six percent relative error in the quantified flow rate, and a four percent relative area in the measured flow area of the flow phantom. The accuracy of these measurements may be further improved by increasing the spatial resolution of the velocity image using a smaller pixel size
110
Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
F l o w R a t e = 100 m L / m i n
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Figure 10. Sixteen of the twenty-four raw Fourier-velocity images of countercurrent flows acquired from the mid-length of the flow phantom. Reprinted with permission from Poh [8]. (the pixel size used in this study was 0.34 mm2). The readers should take note, however, this accuracy does not necessary represent the accuracy in measuring flow in hollow-fiber membrane modules, which depends upon the inner diameter of the hollow-fiber membrane and the spatial resolution of the MR images. 5.2.1.2DFT Technique We acquired twenty-four Fourier-velocity images, each with a velocity bin size of two mm/s, across the mid-length of the flow phantom. Figure 10 shows sixteen of the twenty-four raw Fourier-velocity images acquired. The image artifact was caused by the Gibbs ringing [36-38] and noise [39,40] generated from the receiver coil during image acquisition and was removed during image processing [8]. The percent flow rates in each velocity bin relative to the total flow rate in each tube determined from these Fourier-velocity images were in good agreement with the analytical values derived from the Poiseuille's equation for a fully developed laminar pipe flow for mean velocities ranging from 2 to 14 mm/s (Figure 11). For a fully developed laminar pipe flow, this percent flow rate relative to the total flow rate is predicted from the Poiseuille's 111
Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
FI0w Rate = 100 mLJmin
Re = 135
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equation to increase linearly with increasing mean velocity (Appendix). The discrepancies between the experimental and analytical values at the mean velocities of 16 and 18 mm/s may have been caused by experimental errors such as inhomogeneity in sensitivity profile of the radio frequency coil, finite velocity resolution in the Fourier-velocity images (each image covers a specific velocity bin size, and the velocity resolution decreases with increasing velocity bin size), Gibbs artifact [36-38], noise [39,40] generated from the receiver coil, and variations in flow velocity during the image acquisition period. @
EXAMPLES OF EXPERIMENTAL RESULTS FROM HOLLOWFIBER HEMODIALYZERS
All experiments were performed on a Siemens 1.5T MAGNETOM Vision whole-body MR imager (Siemens AG, Erlangen, Germany) with a maximum gradient strength of 25 mT/m and a rise time of 600 ItS. All image processing was performed using user-written MATLAB programs together with the MATLAB Image Processing Toolbox 2.2 software (The MathWorks Inc., Natick, MA, USA) running on a Linux workstation.
112
Nonintrusive Characterization Of Fluid Transport PhenomenaIn Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
POLYAMIDE HEMODIALYZER Flow Rate: 400 mL/min
Blood Compartment
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Figure 12. Blood-side velocity distribution at the middle cross section of a polyamide hemodialyzer. Reprinted with permission from Poh [8].
We characterized the flow distribution in the blood compartment of hemodialyzers by introducing a three-millimolar cupric sulfate solution, prepared from distilled water, through the blood compartment using a nonpulsatile pump, while the dialysate compartment was filled with the same solution and sealed. Similarly, we characterized the flow distribution in the dialysate compartment by introducing the solution through the dialysate compartment, while the blood compartment was filled with the solution and sealed. In addition, we characterized the flow profile in each compartment and quantified the local ultrafiltration rates by introducing the solution through the blood compartment; we recycled the outflow from the blood compartment to the dialysate compartment countercurrently. The flow rates were between 200 to 1000 mL/min, and five imaging planes were acquired from different transverse cross sections of each hemodialyzer. Detailed experimental methods and image processing methodology have been reported in Poh [8]. Figures 12 and 13 show the velocity distributions in the blood and dialysate compartments of a polyamide hemodialyzer. These velocity images were acquired from the middle cross section (mid-length) of the hemodialyzer using the 2DPC technique with nominal blood and dialysate flow rates of 400 and 600 mL/min, respectively. The polyamide hemodialyzer contains 10,000 hollow fibers, each with an inner diameter of 215 ~m and a wall thickness of 50 113
Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
POLYAMIDE
HEMODIALYZER
Dialysate Compartment
Flow Rate: 600 mL/min
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~tm. It has an ultrafiltration coefficient of 71 mL/Hr/mmHg and a total membrane surface area of 1.7 m 2. Figures 12 and 13 show that the flow distribution in the blood compartment of the polyamide hemodialyzer was uniform and that in the dialysate compartment was nonuniform. Figure 14 shows the dialysate-side axial velocity distributions acquired from the imaging planes right across the dialysate ports of the polyamide hemodialyzer at nominal flow rates between 200 and 1000 mL/min. Figure 14 shows that the axial dialysate-side flow near to the dialysate ports moved preferentially at the top region of the cross section of the hemodialyzer. Figure 15 shows that the percent flow rate increased linearly with increasing mean velocity at the mid-length of a cellulose triacetate hemodialyzer, indicating that the flow in the blood compartment of the hemodialyzer had a fully developed laminar flow profile (Appendix). On the other hand, the flow in the dialysate compartment of the hemodialyzer at the same imaging plane was slightly skewed towards higher mean velocities (positive skewness in mean velocities), indicating that slightly more flow moved at lower mean velocities than at higher mean velocities. The data in Figure 15 were obtained from the Fourier-velocity images reconstructed by the 2DFT technique at a nominal flow rate of 280 mL/min. The cellulose triacetate hemodialyzer contains 12,000 hollow fibers, each with an inner diameter of 200
114
Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
CELLULOSE TRIACETATE HEMODIALYZER Dialysate C o m p a r t m e n t
Dialysate In
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Figure 14. Dialysate-side axial velocity distributions acquired from the imaging planes
right across the dialysate ports of a polyamide hemodialyzer at nominal flow rates between 200 and 1000 mL/min. Reprinted with permission from Poh [8]. lam and a wall thickness of 15 pm. It has an ultrafiltration coefficient of 36 mL/hr/mmHg and a total membrane surface area of 1.9 m 2. Figure 16 shows the local maximum velocities and ultrafiltration rates quantified from a polyethersulfone hemodialyzer using the 2DFT technique at nominal flow rates of 300 and 600 mL/min, respectively. The local maximum velocities and ultrafiltration rates decreased monotonically along the length of the hemodialyzer, indicating that the transmembrane pressure decreased along the length of the hemodialyzer. The polyethersulfone hemodialyzer contains 11,200 hollow fibers, each with an inner diameter of 200 ~tm and a wall thickness of 30 ~tm. It has an ultrafiltration coefficient of 84 mL/hr/mmHg and a total membrane surface area of 1.9 m 2. These experimental results served as an illustration of the capabilities of the 2DPC and 2DFT techniques. Detailed experimental results and discussions have been published elsewhere [8-14].
115
N o n i n t r u s i v e Characterization O f F l u i d Transport P h e n o m e n a In H o l l o w - F i b e r Membrane Modules U s i n g E x p e r i m e n t a l A p p r o a c h - Poh
Flow Rate = 280 mL/min 20
MRI: An Innovative
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LIMITATIONS OF THE 2DPC AND 2DFT TECHNIQUES
Like all other experimental tools, these flow-imaging techniques have their own limitations. Both flow-imaging techniques are only sensitive to flow moving in the flow-encoding direction. In the examples of experimental results presented earlier, the flow-encoding direction was in the axial direction of the hemodialyzer. In reality, the flow in the dialysate compartment of a hemodialyzer moves in all other directions in addition to the axial direction due to the complex flow pathway created by the presence of several thousands of hollow fibers in the hemodialyzer. Moreover, these flow-imaging techniques are only applicable to flow with steady velocity due to their long acquisition times (approximately one minute in the 2DPC and ten minutes in the 2DFT). Any changes in flow velocity during the acquisition period can lead to artifacts in the acquired MR images and subsequently can lead to inaccuracy in the recorded velocities. However, these flow-imaging techniques can be modified by using rapid imaging techniques to account for pulsatile flow seen in all hemodialysis therapies.
116
Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
120
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Local blood-side maximum velocities and ultrafiltration rates along the length of a polyethersulfone hemodialyzer at nominal flow rates of 300 and 600 mL/min. Reprinted with modification from Hardy et al. [14].
It is also worthwhile pointing out that the 2DPC technique is only capable of measuring unidirectional flow in a hollow-fiber membrane module. Any countercurrent velocities contained in a single pixel are averaged over the whole pixel. Moreover, if the flow is oblique instead of perpendicular to the imaging plane, the velocity can be underestimated [34] by a factor of cos(B), where fl is the oblique angle between the flow direction and the imaging plane. For a small r, the error caused by the obliqueness of flow is small since cos(r) is approximately one for a small ft. If fl is known and is smaller than 45 °, the underestimated velocity can be corrected by dividing it by cos(r). In addition to the obliqueness of flow, eddy-current [41 ] induced by the changes in gradient waveforms during data acquisition from the 2DPC technique causes some bias in the velocity determined from the measured phase shift. The eddy current causes a small variation in the phase shift and hence a small variation in the velocity. The eddy-current effect is often seen as a variation in phase shift in the phase images with stationary flow. The eddycurrent effect can be corrected after image acquisition by examining the variation in phase shift of a stationary flow, and it has been corrected in the
117
Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
experimental results shown in this chapter. The correction for the eddy-current effect has been discussed in Walker et al. [41] and Poh [8]. In the 2DPC technique, the measured velocities are also distorted by the partial volume effect [42-45]. The measured velocity in each pixel is determined by the phase of the resultant magnetization, which is affected by all MR-visible protons within the voxel (volume element). In the case of the hemodialyzers, each voxel contains a mixture of protons in three environments: inside the hollow fibers, within the wall of the hollow fibers, and outside the hollow fibers. Different voxels will have different percentages of protons in each environment because of the random distribution of the hollow fibers relative to the imaging grid. In the characterization of the blood-side flow distribution, the fluid inside the hollow fibers was moving, while that outside the hollow fibers was stationary; the opposite occurred in the characterization of the dialysate-side flow distribution. Thus, in each voxel, the signal arises from both stationary and moving protons; the measured velocity in each pixel is therefore distorted and is generally lower than the true velocity. Nevertheless, the region-to-region variation in velocity can be estimated by measuring the mean velocities in different region-of-interest (ROI) large enough to include sufficient number of pixels to average out the variation arising from differences in the number of hollow fibers in each voxel. That is, by determining the mean velocity in one region, we will have derived a quantity that is proportional to the total flow rate of that region. This means velocity can then be used to compare the differences in flow in different ROI of the hemodialyzers. On the other hand, in the 2DFT technique, the measurable velocity resolution is limited by the velocity bin size selected for each Fourier-velocity image. Theoretically, we can improve the measurable velocity resolution by decreasing the velocity bin size. However, there is a tradeoff between velocity resolution, acquisition time, and signal-to-noise ratio for a given range of velocities. For a given range of velocities, a higher velocity resolution causes a lower signal-to-noise ratio and requires a longer acquisition time due to the increase in the number of Fourier-velocity images required for a smaller velocity bin size. 8.
CONCLUSIONS
We introduced two innovative, nonintrusive flow imaging techniques using MRI called the 2DPC and 2DFT techniques with the hope that these flowimaging techniques will provide a greater insight into the fluid transport phenomena in hollow-fiber membrane modules. These flow-imaging techniques provide a valuable nonintrusive experimental tool for studying flow distribution in hollow-fiber membrane modules. The experimental results obtained from these flow-imaging techniques give the actual fluid transport behavior in the 118
Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
hollow-fiber membrane modules that cannot be predicted from numerical models and provide us a better physical understanding of the fluid transport phenomena in these modules. Moreover, an improved understanding of fluid transport phenomena in these modules may lead to the development of more efficient hollow-fiber membrane modules. These flow-imaging techniques can also be used to validate, improve, or develop numerical models capable of predicting separation performance of hollow-fiber membrane modules under different engineering designs and operating conditions. Some examples of experimental results obtained from hollow-fiber hemodialyzers are presented to illustrate the capabilities of these flow-imaging techniques. The readers are encouraged to refer to the references cited throughout this chapter for a full appreciation of these and other MRIbased flow-imaging techniques.
REFERENCES [ 1] [2] [3] [4]
[5] [6] [7] [8] [9] [10] [11] [12] [13]
B.E. Hammer, C.A. Heath, S.C. Mirer, G. Belfort, Quantitative flow measurements in bioreactors by nuclear magnetic resonance imaging, Biotechnology, 8 (1990) 327-30. C.A. Heath, G. Belfort, B.E. Hammer, S.D. Mirer, J.M. Pimbley, MRI and modeling of flow in hollow-fibre bioreactors, AIChE J, 36 (1990) 547-58. C. Donoghue, M. Brideau, P. Newcomer, B. Pangrle, D. BiBiasio, E. Walsh, Use of magnetic resonance imaging to analyze the performance of hollow-fiber bioreactors, Ann N Y Acad Sci, 665 (1992) 285-300. B.J. Pangrle, E.G. Walsh, S. Moore, D. DiBiasio, Investigation of fluid flow patterns in a hollow fiber module using magnetic resonance imaging, Biotech Tech, 3 (1989) 67-72. J. Zhang, D.L. Parker, J.K. Leypoldt, Flow distributions in hollow fiber hemodialyzers using magnetic resonance fourier velocity imaging, ASAIO J, 41 (1995) M678-82. S. Laukemper-Ostendorf, H.D. Lenke, P. Blumler, B. Blumich, NMR imaging of flow in hollow fiber hemodialyzers, J Membr Sci, 138 (1998) 287-95. T. Osuga, T. Obata, H. Ikehira, S. Tanada, Y. Sasaki, H. Naito, Dialysate pressure isobars in a hollow-fiber dialyzer determined from magnetic resonance imaging and numerical simulation of dialysate flow. Artif Organs, 22 (1998) 907-9. C.K. Poh, Characterization of fluid transport in artificial kidneys - A MR/approach, MS Thesis, University of Kentucky, Lexington, KY, USA, 2001. C.K. Poh, P.A. Hardy, W.R. Clark, D. Gao, Characterization of flow in hemodialyzers using MRI, Proc Intl Soc Mag Reson Med, 9 (2001) 1984. C.K. Poh, P.A. Hardy, D. Gao, W.R. Clark, Measurement of local ultrafiltration rate in a hemodialyzer: a novel MRI approach (abstract), ASAIO J, 47 (2001) 160. C.K. Poh, P.A. Hardy, Z. Liao, Z. Huang, D. Gao, W.R. Clark, Characterization of flow distribution in hemodialyzers using MR/(abstract), ASAIO J, 48 (2002) 179. C.K. Poh, P.A. Hardy, Z. Liao, Z. Huang, W.R. Clark, D. Gao, Effect of spacer yams on the dialysate flow distribution of hemodialyzers: A MRI study, ASAIO J, in press. C.K. Poh, P.A. Hardy, Z. Liao, Z. Huang, W.R. Clark, D. Gao, Characterization of flow distribution in hemodialyzers - A MR/approach, in preparation.
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Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
[14]
[151 [16] [17] [18] [19] [20] [21] [22] [23]
[24] [25] [26] [27] [28] [29]
[30] [31]
[32] [331
P.A. Hardy, C.K. Poh, Z. Liao, D. Gao, W.R. Clark, The use of magnetic resonance imaging to measure local ultrafiltration rate in hemodialyzers, J Membr Sci, 204 (2002) 195-205. Abragam, The Principle of Nuclear Magnetism, Oxford University Press Inc., New York, USA, 1983. B.R. Friedman, A.P. Salmon, G. Chaves-Munoz, C.R. Merritt, J.P. Jones, Principles of MR/, McGraw-Hill Co., USA, 1989. P.T. Callaghan, Principles of Nuclear Magnetic Resonance Microscopy, Oxford University Press Inc., New York, USA, 1995. M.A. Brown, R.C. Semelka, MR/: Basic Principles and Applications, 2nd ed., Wiley, John & Sons Inc., USA, 1999. T.J. Mosher, M.B. Smith, A DANTE tagging sequence for the evaluation of translational sample motion, Magn Reson Med, 15 (1990) 334-9. D.D. Blatter, D.L. Parke, R.O. Robinson, Cerebral MR angiography with multiple overlapping thin slab acquisition. Part I. Quantitative analysis of vessel visibility, Radiology, 179 (1991) 805-11. E.M. Haacke, A.S. Smith, W. Lin, J.S. Lewin, D.A. Finelli, J.L. Duerk, Velocity quantification in magnetic resonance imaging, Top Magn Reson Imaging, 3 (1991) 34-49. P. Mansfield, M. Bencsik, Fluid flow in porous systems, Magn Reson Imag, 16 (1998) 451-4. R.B. Miles, W.R. Lempert, Quantitative flow visualization in unseeded flows, Annu Rev Fluid Mech, 29 (1997) 285-326. F.W. Wehrli, Time-of-flight effects in MR imaging of flow, Magn Reson Med, 14 (1990) 187-93. A. Constantinesco, J.J. Mallet, A. Bonmartin, C. Lallot, A. Briguet, Spatial or flow velocity phase encoding gradients in NMR imaging, Magn Reson Imag, 2 (1984) 335340. P.M. Pattany, J.J. Phillips, L.C. Chiu, J.D. Lipcamon, J.L. Duerk, J.M. McNally, S.N. Mohapatra, Motion artefact suppression technique (MAST) for MR imaging, J Comput Assist Tomogr, 11 (1987) 369-377. P.J. Keller, F.W. Wehrli, Gradient moment nulling through the Nth moment. Application of binomial expansion coefficients to gradient amplitudes, J Magn Reson, 78 (1988) 145-149. J.G. Pipe, T.L. Chenevert, A progressive gradient moment nulling design technique, Magn Reson Med, 19 (1991) 175-179. P.R. Moran, A flow velocity zeugmatographic interlace for NMR imaging in human, Magn Reson Imag, 1 (1982) 197-203. F. St~hlberg, B. Nordell, A. Ericsson, T. Greitz, B. Persson, G. Sperber, Quantitative study of flow dependence in NMR images at low flow velocities, J Comput Assist Tomogr, 10 (1986) 1006-1015. G.T. Gullberg, S. Margaret, and F.W. Wehrli, A mathematical model for signal from spins flow during the application of spin echo pulse sequences, Magn Reson Imag, 6 (1988) 437-461. A. Caprihan, and E. Fukushima, Flow measurements by NMR, Physics Reports (Review Section of Physics Letters), 198 (1990) 195-235. J.M. Pope, S. Yao, Quantitative NMR imaging of flow, Concepts in Magnetic Resonance, 5 (1993) 281-302.
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Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
[34]
[351 [36] [37] [38] [39] [40]
[41] [42] [431 [44] [45]
N.J. Pelc, G. Sommer, K.C.P. Li, T.J. Brosnan, R.J. Herfkens, D.R. Enzmann, Quantitative magnetic resonance flow imaging, Magnetic Resonance Quarterly 10 (1994) 125-147. D.C. Wilcox, Basic Fluid Mechanics. DCW Industries Inc., California, USA, 1997, pp. 214. R.M. Henkelman, and M.J. Bronskill, Artifacts in magnetic resonance imaging, Rev Magn Reson Med, 2 (1987) 1-126. D. L. Parker, G.T. Gullberg, P.R. Frederick, Gibbs artifact removal in magnetic resonance imaging, Med Phys, 14 (1987) 640-5. S. Amartur, E.M. Haacke, Modified iterative model based on data extrapolation method to reduce Gibbs ringing, J Magn Reson Imaging, 1 (1991) 307-17. C.N. Chen, V.J. Sank, S.M. Cohen, D.I. Hoult, The field dependence of NMR imaging I. Laboratory assessment of signal-to-noise ratio and power deposition, Magn Reson Med, 3 (1996) 722-9. A. Macovski, Noise in MRI, Magn Reson Med, 36 (1996) 494-7. P.G. Walker, G.B. Cranney, M.B. Scheidegger, G. Waseleski, G.M. Pohost, A.P. Yoganathan, Semiautomated method for noise reduction and background phase error correction in MR phase velocity data, J Magn Reson Imaging, 3 (1993) 521-30. R.L. Wolf, R.L. Ehman, S.J. Riederer, P.J. Rossman, Analysis of systematic and random error in MR volumetric flow measurements, Magn Reson Med, 30 (1993) 8291. C. Tang, D.D. Blatter, D.L. Parker, Accuracy of phase-contrast flow measurements in the presence of partial-volume effects, J Magn Reson Imaging, 3 (1993) 377-85. C. Tang, D.D. Blatter, D.L. Parker, Correction of partial-volume effects in phasecontrast flow measurements, J Magn Reson Imaging, 5 (1995) 175-80. R.M. Hoogeveen, C.J. Bakker, M.A. Viergever, MR phase-contrast flow measurement with limited spatial resolution in small vessels: value of model-based image analysis, Magn Reson Med, 41 (1999) 520-8.
APPENDIX
This appendix shows the derivation of the linear relationship between the percent flow rate in each velocity bin and the mean velocity for a fully developed laminar, circular tube flow from the Poiseuille's equation. From the Poiseuille's equation, we have:
vi
2Q
=~(R n.R4
2 -r/2)
(A.1)
where v~ is the axial velocity at radius, r~ of the tube, Q is the total flow rate in the tube, and R is the inner radius of the tube. Rearranging Eq. (A. 1), we have:
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Nonintrusive Characterization Of Fluid Transport Phenomena In Hollow-Fiber Membrane Modules Using MRI: An Innovative Experimental Approach - Poh
/
V i ;r~ 4
t~ = aiR 2 V
(A.2)
2Q
The concentric flow area, Ai of a laminar tube flow at radii, r~ and Fi+1 is:
Ai --;rg'(ri21--ri 2 )
(h.3)
Substituting Eq. (A.2) into Eq. (A.3) and simplifying, we have:
;rt.2R4 A i "- ~ ( v
20
i
-Vi+l)
(m.4)
From our 2DFT technique, the velocity bin size, ( v i - Vi+l) is a constant. Therefore, A; is independent of (vg- V;+l) and is a constant for a given R and Q from Eq. (A.4). The flow rate, Q; at a given concentric flow area, A; is given by:
Qi =(v)iAi
(A.5)
where
; is the mean velocity corresponding to A; and is given by: 1
(v)i = --~(vi + vi+, )
(A.6)
Eq. (A.5) shows that Qi is linearly proportional to ;. The percent flow rate, %Q; in each velocity bin, (v;- V~+l) is determined by dividing Eq. (A.5) with Q and then multiplying by 100 percent:
=~
Q
x 100%
(A.7)
Therefore, from Eq. (A.7), the percent flow rate in each velocity bin, %Q is linearly proportional to the mean velocity, ; where A~ is a constant for our 2DFT technique, and Q is a constant for a fully developed laminar, circular tube flow.
122
Membrane Contactors and Environmental Applications
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New Insights into Membrane Science and Technology: Polymeric and Biofunctional Membranes D. Bhattacharyya and D.A. Butterfield (Editors) © 2003 Elsevier Science B.V. All rights reserved.
Chapter 6
Industrial applications and opportunities for membrane contactors R. Klaassen, P.H.M. Feron, R. van der Vaart, A.E. Jansen
TNO Institute of Environmental Sciences, Energy Research and Process Innovation, Department Chemical Engineering PO Box 342, 7300 AH Apeldoorn, The Netherlands phone: +31 55 493196 fax: +31 55 5493410 email: [email protected] SUMMARY
In a membrane contactor membrane separation is fully integrated with another separation technology, like extraction or absorption, in order to exploit the benefits of both technologies fully. Membrane contactor applications can be found both in water and gas treatment. Several recently developed applications that have been introduced in industry - aromatics recovery from a process water, selective removal of heavy metals from a galvanic process bath and ammonia recovery from an off gas s t r e a m - and new applications under developmentCO2 membrane gas absorption and membrane stripping- will be discussed technically and economically to show the wide scope of the technology. 1.
INTODUCTION
Stringent demands are put onto new separation technologies due to stricter product quality requirements, environmental legislation, and energy efficiency demands and needs for cost reduction. In order to meet these ever increasing needs there is a tendency to combine processes to a hybrid process. In a membrane contactor a membrane separation is not only combined with an extraction or absorption process but both processes are fully integrated and incorporated into one piece of equipment. In this way advantages of both processes can be fully exploited.The membrane offers a flexible modular energy efficient device with a high specific surface area. The extraction or absorption process can offer a very high selectivity and a high driving force for transport even at very low concentrations. Applications of membrane contactors can be found in gas and water
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Industrial Applications And Opportunities For Membrane Contactors - Klaassen
treatment. In the membrane gas absorption process the gas stream to be treated is brought very efficiently into contact with an absorption liquid in the membrane contactor. Typical examples are CO2 removal from flue gas or indoor air, flue gas desulphurisation, indoor air conditioning, ammonia recovery from off gasses or mercury removal from natural gas. In the pertraction process water is treated in a membrane contactor with an extraction liquid. Typical applications are hydrocarbons (e.g. aromatics, phenol, haloginated hydrocarbons) removal and recovery from waste and process water and selective recovery of heavy metals from e.g. galvanic bath liquids or waste water. In a membrane contactor also ammonia can be removed and recovered from waste water in a combined stripping and absorption process. Many different developments in the field of membrane contactor technology are going on worldwide. Development of a specific application normally starts with a proof of principle followed by a feasibility study, further development, pilot plant tests and finally full-scale demonstration on site. After a general introduction on membrane contactors some examples of successful developed full-scale industrial membrane contactor installations will be discussed in this paper. These industrial reference installations help to gain confidence of end users in the membrane contactor technology. Thus further market introduction will be facilitated as well as development of new applications for membrane contactors. 2.
MEMBRANE GAS ABSORPTION
2.1
Principle
Membrane gas absorption is a gas-liquid contacting operation [ 1,2]. The key element in the process is a microporous hydrophobic hollow fibre membrane. The process is illustrated in figure 1 for removal of component X from a gas stream. The gas stream is fed along one side of the membrane where an absorption liquid is flowing at the other side of the membrane. The hydrophobic membrane wall keeps gas phase and absorption liquid separated from each other. The absorption liquid is chosen in such a way that it has a high affinity for component X. Component X will now diffuse through the gas filled pores of the membrane to the other side of the membrane where it is absorbed in a liquid phase. Absorption in the liquid phase takes place either by physical absorption or by a chemical reaction. This determines the selectivity of the process. The membrane used gives no contribution to the selectivity: the membrane's role is to keep the two phases separated and to provide a large contacting surface area.
126
Industrial Applications And Opportunities For Membrane Contactors - Klaassen
Hollow fiber
gas
=,.,ou,v,u,, liquid ~
gas
Membrane
Figure 1: Membrane gas absorption For operation of membrane gas absorbers it is essential that liquid phase and gaseous phase do not mix. This means that the absorption liquid is not allowed to enter the pores. This is influenced by pore size, pressure difference across the membrane and the interaction of the absorption liquid with the membrane material. This is described by the Laplace equation: d P = - 2 (y/r) . c o s (~o)
( 1)
where: d P - pressure difference between liquid and gas [Pa] 7 - surface tension of the absorption liquid [N/m] r - pore radius [m] (p - contact angle for liquid and membrane material The membrane pores will not be wetted if the contact angle is greater than 90 ° and the pressure difference is limited for a given pore size. Suitable membrane materials for aqueous absorption liquids are non-polar polymers such as polypropylene, polyethylene and Teflon. 2.2
Membrane module The design of a membrane absorber cannot be based on conventional
127
Industrial Applications And Opportunities For Membrane Contactors - Klaassen
available hollow fibre membrane modules. These module designs have historically been based on the use of filtration duties and consist of a bundle of hollow fibres in a tubular housing. This gives ill-defined flow conditions on the outside of the fibre bundle. When membranes are used for contacting duties it is necessary that flow conditions are well def'med on both sides of the membrane to achieve good mass transfer. Furthermore it is essential that, given the often large volume flows of flue or off gas streams, the module design gives low pressure drops and is easy to scale up. TNO has patented a new membrane module design to overcome the problems with conventional membrane modules. Figure 2 shows this new membrane module design [3] for contacting duties. The module consists of a rectangular housing with fibres at well-def'lned positions. The gas flow is perpendicular to the fibres. The absorption liquid is fed through the fibres. Elements of the module can be coupled parallel and in series. In this way the module can be easily adjusted to the flow rate of the gas stream to be treated and the desired removal efficiency.
inlet
outlet
inlet
Figure 2: New membrane module
Important features of the module are: well defined flow conditions on both sides of the membrane, easy scale up, high mass transfer, low pressure drop and high specific surface area. Specific surface area's of well over 1000 rn2/m3 can be reached. 2.3
Membrane gas absorption versus conventional contactors
Conventional G-L contactors, such as packed towers, preferably operate
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Industrial Applications And Opportunities For Membrane Contactors - Klaassen
with relatively large gas flows under steady-state conditions. Membrane gas absorption can be an eminently suitable alternative for 'smaller' gas flows and operation at variable loads. Especially by using hollow fibre membranes it is possible to develop a very compact gas-liquid contactor. Hollow fibre membrane modules combine a high specific surface area, a well-def'med gas flow regime and short diffusion ways in a modular piece of equipment. The membrane gas absorption technology is able to treat gas streams with extreme variations in gas flow and or concentrations of the components to be removed. This is possible at low costs in very compact equipment. The reduction in volume and weight compared to a conventional absorber can be more than a factor 10. Through the use of hollow fibre membranes the membrane gas absorption process can offer operational and economical advantages over conventional spray towers or packed columns. Operational advantages include independent Gas/Liquid control, flexible operation, optimal load of the absorption liquid, no entrainment, flooding or foaming and modular low weight and very compact equipment. Economical advantages include low investment costs, low pumping power for the absorption liquid and the fact that no expensive civil engineering work is required. That gas absorption membranes can result in very compact equipment can be illustrated by the following example: For a coal fired power plant of 645 MWe a design has been made for a conventional absorber (spray tower) and a membrane gas absorption unit. A flue gas flow of 2"106 nm3/hr has to be treated in this case. The calculated overall volumes of the absorbers are given below. Conventional SO2 absorber Calculated membrane gas absorption
2.4
- volume 9000 m 3 - volume 250 m 3
Applications
Membrane gas absorption can be applied for removal of those components from a gas stream where a suitable absorption liquid is available. A suitable absorption liquid has a high affinity for the components to be removed and does not wet the membrane. The majority of absorbents used in conventional gas absorption processes can also be used for membrane gas absorption. Applications [4,5] for the use of membrane gas absorption are listed below. Flue-gas and off-gas containing e.g. SO2, HC1, NH3 or HES can be treated to meet emission standards in many industrial situations. Recovery and re-use of CO2 from flue gas, biogas and off-gases in horticultural industry, beverage production and other industrial
129
Industrial Applications And Opportunities For Membrane Contactors - Klaassen
applications. Upgrading and desulphurisation of biogas from anaerobic digesters and landfills. Acid gas removal and dehydration of natural gas. Acid gas removal from fuel gas mixtures. Mercury removal from natural gas, flue gas or glycol overheads. Alkene-alkane separation in petrochemical industry. Indoor air treatment for e.g. tobacco smoke components. Due to the modular equipment used in membrane processes application of gas absorption membranes is also possible on a small scale. This is particularly valuable for applications in the field of environmental technology, as it enables an efficient and economical treatment of small-scale emission sources. If regeneration of the absorbent cannot be done one site, central regeneration is another possibility. In paragraph 2.5 the recovery of ammonia with membrane gas absorption is discussed. In paragraph 2.6 developments with respect to CO2 absorption are shown. Another industrial application of membrane gas absorption is the transfer of unstable C102 to potable water [27]. Also the recovery of VOC's from a paint booth exhaust is recently shown under minipilotplant conditions [28].
2.5.
Membrane gas absorption installation for ammonia recovery An industrial membrane gas absorption unit for ammonia recovery has been installed at a dyes intermediates production plant of Aliachem in Pardubice the Czech Republic (Figure 3). In this plant dyes intermediates are produced batch wise in pressurised reactors. Ammonia is used as a reactant in the reactor. Depressurisation of the reactor results in an ammonia containing off gas stream, which gives a significant loss of ammonia. In the membrane gas absorption unit the total off gas stream is treated with water as absorption liquid. The ammonia is recovered as an aqueous ammonia solution of > 20 wt% which can be re-used in the dyes intermediates production process. The membrane gas absorption unit is successfully in operation since June 1999; aqueous ammonia solutions of 27 wt% ammonia have been produced and the emission of ammonia to the environment is reduced with 99.9%. The installation has a capacity to absorb 50 kg/hr ammonia and has proved to be very easy in operation. Variation in gas flow rate or ammonia concentration can be handled without any problems.
130
Industrial Applications And Opportunities For Membrane Contactors - Klaassen
Figure 3: MGA unit on site in Pardubice, Czech Republic 2.6
C02 removal
2.6.1 Introduction CO2 membrane gas absorption has been under development at TNO for a number of years. The potential of this technology has been identified in a number of assessment studies [6,7]. The ensuing customer-driven development has been focusing on the production of carbon dioxide from flue gas with the aim of supplying high purity carbon dioxide to the horticultural industry to increase crop yield [8,9]. Other commercial applications are in the field of life support (spacecrafts, submarines, operating theatres, consumer products, biogas treatment). In recent years there has also been an increased interest in CO2-removal and subsequent storage in view of the enhanced greenhouse effect. Figure 4 shows an overview of the atmospheric pressure applications currently being investigated. The process is based on novel membrane absorption method patented by TNO [10], which employs dedicated aqueous absorption liquids, called CORAL, based on amino-acids and alkaline salts. Benefits of these liquids over conventional amines are: - Stable operation with cheapest membranes available viz. polyolefin membranes. - Oxygen stability. - Improved corrosion resistance. - No vapour pressure, hence no loss of absorption liquids through evaporation.
131
Industrial Applications And Opportunities For Membrane Contactors - Klaassen
CO~-MGA application range 100 --0-
x "
~
• O~ (J
~
--
Greenhouse Blogas Submarine Space Medical Consumer
0.1 --o--
0.01 0.001
1
1000
Capacity [kg C021h ]
Figure 4:CO2 membrane gas absorption applications (atmospheric pressure) currently under development.
2.6.2 C02-production for greenhouses In the Netherlands carbon dioxide is used to promote plant growth in greenhouses. Depending on the crop, the production can be increased by 25% by increasing the CO2-concentration to 500 ppm. In temperate climates it is necessary to heat these greenhouses, which makes the horticultural industry in the Netherlands a large consumer of natural gas. Future heat supply systems will be largely based on cogeneration plants (gas engines and gas turbines) as significant energy savings can be achieved. Energy savings can be increased if the heat supply is coupled to carbon dioxide supply. This will incite growers to buy heat from the cogeneration plant rather than to use their own boilers. These boilers are needed to cover peak demand and emergencies. Also the amount of available carbon dioxide per unit heat will be larger in case of combined production of heat and electricity, thus allowing for an improved production. The heat demand and CO2 demand are generally anti-cyclic. During periods of high heat demand (during winter and at night) little or no CO2 is needed to maintain the desired CO2 levels in the greenhouse. During periods of high CO2 demand (during summer and during the day) little or no heat is needed. The thermal energy of the cogeneration plant can either be used for greenhouse heating or for the production of carbon dioxide. This will also extend the operating time of the cogeneration plant to daytime hours, which is an additional benefit in terms of capacity credits. A schematic drawing of a cogeneration plant supplying heat and carbon dioxide to greenhouses is shown in Figure 5.
132
IndustrialApplicationsAndOpportunitiesForMembraneContactors- Klaassen
Greenhouses Flue gas
I
ico2 (loo°/o) ,
l
H.t
N,ura,
Electricity 1
T Natural gas
Grid
Figure 5: Cogeneration plant supplying heat and carbon dioxide to greenhouses
2.6.3 Pilot plant Pilot plant experiments were carried out to achieve the following objectives: - Assessment of mass transfer under a variety of conditions with several liquids, - Assessment of heat consumption, - Assessment of CO2-quality. During the feasibility study a pilot plant has been designed and built for testing in the lab as well as for testing on location with a real flue gas. The basic specifications of the pilot plant are shown in table 1. Table 1: Basic specifications of COz-removal pilot plant CO2-concentration 3.5 % Gas flow rate 25 m3/hr CO2-removal 80% Liquid flow rate 20-1001/hr The pilot plant flow sheet is shown in Figure 6. In case of the lab experiments CO2 and air are mixed via mass flow controllers and fed to a prototype cross-flow membrane absorber (DAM-module). In this absorber is CO2 is transferred from the gas stream to a liquid stream in a counter current fashion. On location a small stream has been taken from the chimney, cooled down and pumped directly to the membrane absorber. The rich liquid is fed to the top of a stripper via a plate heat exchanger. The stripper temperature is maintained at a temperature of around 120°C by an electric heating element. As a result of the temperature increase of the absorption liquid CO2 is liberated from the solution. Water is recovered in the condensor and trickles back into the stripper. The lean liquid exits the stripper at the bottom and is fed back to the membrane absorber via the heat exchanger. Temperatures, pressures, dew
133
IndustrialApplicationsAndOpportunitiesForMembraneContactors- Klaassen
points, flow rates and heat consumption have been measured and are used for the performance assessment. C02
Steam Gas in
Air r v
~Mi i
mo Ju
Rich I exchanger I =' I Heat =I
I "~
! Stripper
Lean
Cooler
~
Heatero
Gas out
.= Gas analyser
Figure 6: Pilot plant flow sheet
The specifications of the prototype cross-flow membrane module designed according to [3] are shown in Table 2. Table 2
Specifications prototype cross-flow membrane module Membrane area 5.0 m 2 Fibre material Polypropylene (Accurel® Q3/2) Nominal pore size 0.3 i.fin Fibre diameter (inside/outside) 0.6/1.0 mm
The pilot plant as operated on location is shown in Figure 7.
Figure 7: Pilot plant on location.
134
Industrial Applications And Opportunities For Membrane Contactors - Klaassen
2.6.4
Mass transfer performance
An important performance parameter is the mass transfer in the membrane gas absorption process, as this will determine the membrane area requirement. Membrane area is expected to represent the largest investment component. The mass transfer, incorporating a chemical reaction, is influenced by: - The type of absorption liquid. The reaction kinetics and reaction path are dependent on the nature of the active components and their concentration levels. - The absorption temperature. Higher temperature leads to faster reaction kinetics but the thermodynamic equilibrium will be less favourable. A trade-offhas to be made here. - Fluid dynamics. The fluid dynamic conditions (gas and liquid) might influence the mass transfer to some extent. The effect is however small. - Membrane type. The membrane type should be sufficiently porous not to impose any additional mass transfer barriers. - CO2-10ading in liquid and CO2-concentration levels in gas. If the liquid already has a considerable CO2-content this will limit the mass transfer. A proper regeneration process can achieve this, but a trade-off with the energy consumption has to be made. Also at high CO2-concentration levels mass transfer will be limited because local saturation of the active components might OCCUr.
The mass transfer performance, expressed as a carbon dioxide flux through the membrane, is shown in figure 8 for the CORAL20 liquid at room temperature as a function of the average carbon dioxide partial pressure. The data have been gathered from a variety of experiments carry out with a small pilot plant including regeneration of the absorption liquid. It clearly shows that over a wide range of the carbon dioxide partial pressure the membrane fluxes will increase with increasing partial pressure. A feasibility study has demonstrated that the CO2 can be produced economically from flue gas on a large scale [11]. The economical analysis in this study indicates that at a production capacity of 10 tonnes CO2 per hour the productions costs will be around 50 US$ per tonne CO2. At this cost level the process is cheaper than e.g. CO2 delivered by truck or current CO2-production processes. A development project is currently underway in which a pilot plant producing around 150 kg CO2 per hour will be built. Successful completion of the development will then pave the way towards a demonstration on a large scale (10 tonnes CO2 per hour) for a greenhouse.
135
Industrial Applications And Opportunities For Membrane Contactors - Klaassen
800
700
600
500
.~ 400 300
200
100
0
v
,
'
re [kPal
Figure 8: Measured CO2 membrane mass flux for CORAL20 at room temperature as a function of the logarithmic mean CO2 partial pressure. 3.
PERTRACTION
3.1
Principle
Pertraction is a non-dispersive membrane based liquid-liquid extraction process. Process or wastewater flows contaminated by organic substances such as aromatics or chlorinated hydrocarbons can be cleaned by pertraction. The pollutants are removed from the water by extraction into an organic extractant, which is immiscible with water. The membrane itself has no selectivity. The process selectivity must come from the extractant. The interface between wastewater and extractant is immobilized using a hydrophobic microporous hollow fiber membrane by means of a small transmembrane pressure gradient (-0.1 bar)[ 12,13], see figure 9. Contrary to conventional extraction processes there is no direct mixing of extractant in the wastewater stream. This offers important advantages over conventional extraction: The often difficult and time-consuming separation between water phase and extractant is not necessary. Flows of waste water and extractant are flexible, and can be adjusted
136
Industrial Applications And Opportunities For Membrane Contactors - Klaassen
independently of another, making process optimisation simple and allowing a highly efficient contact between a large volume of (waste) water and a very small quantity of extractant. Pertraction installation can be of very compact construction, thanks to the high specific surface area and good mass transfer of the pertraction membrane modules. No differences in density between the water phase and the organic extractant are required. extractant hollow fibre /^\
J '!
waste water
/
i
'
i
:
!
' i
i
\
/
\
/
\
! s s
o,
i! i! i
~
waste water
t
, l i
!
i
i
'
t
~r
T
extractant
extractant
Figure 9: Principle of pertraction
An advantage of pertraction over stripping and carbon adsorption is that both volatile and non-volatile components can be extracted in one step. In the pertraction process pollutants are concentrated in an extractant (Figure 10). The latter can be either directly re-used or regenerated in e.g. a vacuum film evaporator, and can thus continue to circulate in the system, while the pollutants are released in pure form. pertraction module P(P(P~,.-r., cleaned
" "
contaminated water
... _ ~
~
"
water
~<
loaded extractant
film evaporator regenerated extractant
Figure 10: Pertractie-regeneration process, with pollutants released in pure form.
137
Industrial Applications And Opportunities For Membrane Contactors - Klaassen
The choice of the proper extractant is an essential step in development of new pertraction applications. In -
the selection procedure requirements for the extractant are: A high affinity for the components to be removed from the waste water. Very low solubility in water. Non-toxic. Low viscosity. Possibilities for re-use or regeneration.
Pertraction is especially suitable for removal of hydrophobic organic compounds like halogenated, aromatic and aliphatic organics from aqueous streams. Water streams containing these hydrocarbons can be found at many places: Effluents/process streams from pharmaceutical-, off shore-, automotive-, paint-, petrochemical- and chemical-plants. Water originating from tank cleaning and tank storage firms. Effluents from garages dry cleaning and wood preservation firms. Contaminated ground water. Landfill leachate from sites containing chemical wastes. 3.2
Economics
of Pertraction
A comparative cost model [ 14] has been developed for pertraction and the conventional technologies air stripping and activated carbon treatment. The cost models for all the technologies are based on the same assumptions to obtain comparable results. The results of the cost models are total treatment costs per cubic meter water treated. These costs include depreciation of the installation, energy costs, maintenance and operation. In table 3 results of the cost calculations are shown for trichloroethene based on a waterflow of 10 m3/hr with an inlet concentration of 10.000 ~tg/1. The outlet concentration of trichloroethene is a variable and lies in the range between 1000 ~tg/1 and 1 ~tg/1. The results show that in all the cases calculated pertraction is cheaper than the alternatives (air stripping or activated carbon filtration). Other substances, which can be economically treated by pertraction, include chlorinated solvents (e.g. carbon tetrachloride, chloroform, tetrachloroethene, trichloroethene), PCB's, di- and tri-chlorobenzene, pesticides and higher polycyclic hydrocarbons.
138
Industrial Applications And Opportunities For Membrane Contactors - Klaassen
Table 3: Cost evaluation: trichloroethene. flow rate: 10 m3/hr trichloroethene inlet concentration 10.000 ,g/1 u'-' Outlet concentration
Total costs in USD per m3 treated water for the different selected processes Air stripping; with air treatment
Activated carbon filtration
Pertraction
1000 ~tg/1
USD 0,53
USD 0,83
USD 0,28
100 ~xg/1
USD 0,53
USD 0,89
USD 0,36
10 ~tg/1
USD 0,60
USD 0,92
USD 0,44
l~g~
USD 0,60
USD 0,94
USD 0,51
3.3 Pertraction installation for recovery of aromatic compounds from waste water
A full-scale pertraction installation has been successfully put into operation in 1998 at the chemical firm KoSa Netherlands BV in Vlissingen the Netherlands. The pertraction installation is equipped with Liqui-Cel ® Membrane Contactors [29]. It treats 15 m3/hr waste water originating from a chemical reactor (Figure 1 I). The water is heavily polluted with the product formed in the reactor. This product is an aromatic compounds. In the pertraction installation the organic compound is extracted from the waste water with a feedstock for the reactor as extractant. In this way the water is not only cleaned but also lost product is recycled back to the reactor. Next to this the existing wastewater incineration installation could be put out of operation thanks to the pertraction installation. In this way KoSa saves 5-million m 3 natural gas annually. The unit is more then three years in operation now and is a critical element in the production process.
Figure 11" Industrial pertraction installation.
139
Industrial Applications And Opportunities For Membrane Contactors - Klaassen
4.
EMULSION PERTRACTION
4.1
Principle
Emulsion pertraction [15] is a process evolved from a combination of permeation through membranes and solvent extraction. In the emulsion pertraction process the water phase is kept apart from the emulsion phase by a hydrophobic microporous membrane. The emulsion phase consists of an organic solvent with a dissolved extractant as continuous phase with aqueous droplets of strip liquid dispersed in it. The contact surface between water phase and emulsion phase lies in the pores of the membrane. The metal to be removed from the water phase is bound by the extractant present in the pores of the membrane. At the other side of the membrane the strip liquid continuously regenerates the extractant. The strip liquid is dispersed in the organic phase. There is no need to make the dispersion very stable because the hydrophobic nature of the membrane keeps the waste water and aqueous strip liquid always separated. Therefore no aqueous phase can pass the membrane. As shown in figure 12, the metal indicated as Me is bound to extractant L to form a complex Me-L which is regenerated at the interface of the droplets strip liquid to form for example a concentrated metal sulphate solution. There is no selectivity needed of the membrane itself. The selectivity of the process is provided by the extractant. Hollow fiber
"~
i ........... '-' ...... 'i
i
i
i
~,~j =
/
!
,,_,_27,,__,,_.,_,:
..,.,,-.,.
I I I
Membrane
>
t t '
I
!
-
~
t
t
I
I
~
water
~
emulsion
Me
Metal
L
Extractant
I
water
Figure 12: Principle emulsionpertraction The emulsion used in emulsion pertraction is analogue to the emulsion used in the Liquid Membrane Process (LMP) developed by the university of Graz [16]. The main advantage of pertraction over the LMP process is that the emulsion may be unstable while LMP requires a stabilised emulsion that is difficult to break. In the Stabilized Liquid Membrane (SLM) process only the pores of the 140
Industrial Applications And Opportunities For Membrane Contactors - Klaassen
membrane are filled with the organic extractant. This process suffers from stability problems due to depletion of the extractant phase, which only has the very small volume of the membrane pores. These problems are over come with the emulsion pertraction process where the extractant that is flowing through the hollow fiber membranes can easily make up for any small losses that might occur from the membrane pores. These losses can be fatal to the SLM process. Advantages of emulsion pertraction over normal extraction are: - Extraction and stripping can be carried out in one operation. - No possibility of emulsion formation in the water phase. - Volume of extractant is relatively small. - Process parameters are very flexible. - Phase separation is not necessary. - Large specific surface area because of hollow fibre membranes. - Short diffusion ways. - Compact modular equipment. - Low energy consumption. Emulsion pertraction is suitable for selective removal of heavy metals like copper, zinc, iron, chromium, cadmium or lead from aqueous solutions. The heavy metals can be recovered in a concentrated solution. Lifetime extension of passivating bath liquids is an industrial application for emulsion pertraction. This application is discussed in paragraph 4.2. The principle of emulsion pertraction is also used by Commodore Separation Technologies for the recovery of Cr(VI) from waste water [30].
4.2 Emulsion pertraction installation for life time extension of passivating bath liquids An emulsion pertraction installation (Figure 13) has been tested on site at Rogal in Enschede the Netherlands to extend the lifetime of their passivating bath. Passivating is an operation frequently used in galvanic industry to improve corrosion resistance of galvanized objects. The lifetime of the passivating bath liquid is limited due to accumulation of zinc and iron. After the zinc and iron concentration reach a certain limit the colour and corrosion resistance of the passivated objects falls no longer within the specs and the bath liquid has to be replaced. Selective removal of iron and zinc from the passivating bath with an emulsion pertraction installation can extend the lifetime of the bath. Care must be taken that no of the active elements chromium or cobalt are being removed. Tests on site have shown that the emulsion pertraction installation is easy
141
Industrial Applications And Opportunities For Membrane Contactors - Klaassen
to operate. Operation takes no more then 20 minutes a week. Zinc and iron(III) are removed from the passivating bath with a high selectivity; there is a
Figure 13: Emulsion pertraction installation at site of Rogal
negligible loss of chromium and cobalt. The extractant used is continuously regenerated with sulphuric acid. The membranes are resistant to the on-site conditions and transport rates are sufficiently high that a compact installation can be built. Due to a reduction in the use of raw materials (flesh bath liquid) and less metal sludge to be disposed of the installation gives annual benefits of USD 10,000 whereas the total annual operational costs are only USD 6,000. 5.
M E M B R A N E STRIPPING
5.1
Introduction
The removal of gases, especially oxygen, from water is important in several industrial sectors, e.g., in the microelectronics industry (rinse water) [17], in the electricity production sector (boiler feed water) [18, 19], in the brewing of beer (heavy brewing) [20, 21] and in the oil & gas industry (injection water) [22]. In the above examples, the oxygen needs to be removed to prevent oxidation of substrates, equipment or dissolved components or to prevent biological growth. Several methods are commercially available for oxygen removal [23]. Classical methods involve the (thermal) vacuum stripping of water in a voluminous packed bed, often combined with the application of a stripping gas. Also, the addition of chemicals like hydrazine or sulphite to the water system is in use [24]. Drawbacks of these methods are high energy consumption, the harsh chemical nature of hydrazine, the increase of chemicals concentration (conductivity) in the product-water in the case of sulphite addition, and the 142
Industrial Applications And Opportunities For Membrane Contactors - Klaassen
consumption of chemicals. More sophisticated methods are based on reduction of oxygen with hydrogen over a Palladium catalyst or stripping via a membrane contactor. These methods are intrinsically clean. A drawback of the hydrogen method is the difficult mixing/dissolving of hydrogen in water, requiring voluminous equipment [ 19]. 5.2
Catalytic membrane contactor The use of membrane contactors to improve the dissolution of hydrogen has been described. The catalytic phase may be present on ion exchange beads in between hollow membrane fibres [25] or in a separate subsequent contactor. Drawbacks of these methods are the use of expensive Pd-loaded ion exchangers and the fact that the product water contains excess hydrogen gas. Recently, the proof-of-principle of a fully integrated catalytic membrane contactor (CMC) has been reported [26]. The CMC system is based on the integration of the catalyst phase with a hydrophobic microporous membrane that keeps the hydrogen and water separated. The catalyst is present at the interface where the reaction takes place on the membrane external surface (Figure 14). This system minimises mass transfer resistances, resulting in optimal performance and minimal required membrane surface area, amount of catalyst and amount of hydrogen lost to the water. No pre-mixing of hydrogen with the water stream is required, which results in a much more compact installation. Moreover, this type of membrane is based on commercially available and relatively low-cost hollow fibre membranes. The CMC was evaluated for use in the production of ultrapure water for the semiconductor industry. It is expected that the CMC should remove oxygen at 23-34% lower investment costs, 20-38% lower fixed costs and 35-75% lower operating costs.
Figure 14: Back-scatterelectron microscope image of a Palladium-coatedporous hydrophobic membrane fibre. The highlighted area indicates the Palladium coating. 6. CONCLUSIONS 143
Industrial Applications And Opportunities For Membrane Contactors - Klaassen
Membrane contactors combine properties like modular concept with high specific surface area from membrane technology with the selectivity from conventional separation processes like extraction or absorption. This creates a new generation of highly efficient, compact and flexible separation technologies. Applications for membrane contactors can be found in gas-liquid and liquidliquid contacting operations. Examples are removal of specific compounds from gas streams (e.g. NH3, CO2, SO2, H2S, H20) and selective recovery o f hydrocarbons and heavy metals from aqueous streams. The first full scale industrial membrane contactor installations are in operation in industry now. The success o f these applications depends on the unique advantages o f membrane contactors (flexible capacity, no phase separation, low volume, low weight, extreme phase ratios etc.). On the other hand it is also very important to demonstrate the reliability on the long run under industrial conditions. This will give more confidence in the m e m b r a n e contactor technology for other end users and can lead to a whole new world o f membrane contactor applications.
REFERENCES [1]
[2] [3] [4]
[5] [6]
[7] [8] [9] [10] [11]
Zhang, Q., Cussler, E.L. - Microporous hollow fibres for gas absorption I. Mass transfer in the liquid, J. Membr. Sci., 23(1985) 321-332. Klaassen, R. et. al, - SO2 removal from flue gas with gas absorption membranes, Proceedings Large Chemical plants, p 75-77, October 1992 Antwerpen Belgium. Meulen, B. Ph. ter - Transfer device for the transfer of matter and/or heat from one medium flow to another medium flow., Eur patent 0509031, US patent 5230796. Klaassen, R., Feron, P.H.M., Jansen, A.E. - Membrane gas absorption environmental applications, Proceedings ICOM 96 p 916-917 Feron, P.H.M., JansenA.E., Klaassen, R. - Membrane gas absorption: a new membrane technology for gas separation and purification 6th Aachener Membran Kolloquium 3-5 March 1997 Feron, P.H.M., JansenA.E., Klaassen, R. - Membrane Technology in Carbon Dioxide Removal, Energy Convers. Mgmt., Vol. 33, No. 5-8, pp. 421-428, 1992 Feron, P.H.M - Membranes for carbon dioxide recovery from power plants, "Carbon Dioxide Chemistry: Environmental Issues",, Ch. 24, pp. 236-249 eds. J. Paul, C.M. Pradier, The Royal Society of Chemistry, Cambridge, 1994 Feron, P.H.M., JansenA.E. - Capture of carbon dioxide using membrane gas absorption and reuse in the horticultural industry, Energy Convers. Mgmt. 36(411-414) 1995 Feron, P.H.M., JansenA.E . - The production of carbon dioxide from flue gas by membrane gas absorption, Energy Convers. Mgmt. 38($93-$98)1997 Feron, P.H.M., et al. - Method for gas absorption across a membrane, USP 5749941, EP 0751815 Feron, P.H.M., JansenA.E. - Techno-economic assessment of membrane gas absorption
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Industrial Applications And Opportunities For Membrane Contactors - Klaassen
[12] [13]
[14] [15]
[16] [17]
[18] [19] [20] [21] [22] [23] [24]
[25] [26]
[27]
[28]
[29] [30]
for the production of carbon dioxide from flue gas, Greenhouse Gas Control Technologies, pp. 53-58 Pergamon, 1999 Kiani, A., Prasad, R., Bhave, R.R., Sirkar, K.K. - Solvent extraction with immobilized interfaces in a microporous hydrophobic membrane, J. Membr. Sci., 20(1984) p125-145 Klaassen, R., Jansen, A.E., Bult, B.A., Oesterholt, F.I.H.M., Schneider, J., - Removal of hydrocarbons from waste water by pertraction - Proceedings 7th International Symposium on Synthetic Membranes in Science and Industry. Ttibingen Germany August 29- September 1, 1994 p 316 - 319 Klaassen, R., Jansen, A.E., Bult, B.A., Oesterholt, F.I.H.M., - Pertraction of hydrocarbons from waste water streams; Final report, ref. nr. 94-039, March 1994. Klaassen, R., Jansen, A.E., - Selective removal of heavy metals with emulsion pertraction, Minerals, Metals and the Environment II p207-216 The Institution of Mining and Metallurgy London 1996 Draxler, J., Marr, R. - Emulsion liquid membranes, Chem.Eng.Process., 20 (1986) 319329 Yagi, Y., Y. Kamasa, Ohmi, T. - Advanced ultrapure water systems with low dissolved oxygen for native oxide-free wafer processing, IEEE Transactions on semiconductor manufacturing, 1992.5: p. 121-127. d'Angelo, P., - Oxygen removal: theory and potential use of deoxygenation membranes in the utility industry, Ultrapurewater, 1995. July/Aug: p. 60-63. Matt, K., - Preventing corrosion by catalytic reduction of oxygen in boiler and recirculation water, Reports Sci. Technol., 1997.59: p. 52-57. Spielberger, F., - Neue Wege der Sauerstoff entfemung im brauwasser. Brauwelt, 1991.23: p. 975-979. Kohnke, H.-V., - Wasserentgassung und Einsatz entgasten Processwassers. Brauwelt, 1991: p. 1357-1359. Henriksen, N., - Deaeration of water, 1985. Butler, I.B., M.A.A. Schoonen, Rickard, D.T., - Removal of dissolved water from water:a comparison of four common techniques, Talanta, 1994.41: p. 211-215. Chabria, M.C., Sharma, M.M., - Kinetics of absorption of oxygen in aquous solutions: (1) acidic chromous chloride, (2) acidic titanous chloride and (3) ammoniacal cuprous chloride. - Chem. Eng. Sci., 1974.29: p. 993-1002. Li, K., I. Chua, Teo, W.K., - Removal of dissolved oxygen in ultrapure water production using a membrane reactor, Chem. Eng. Sci., 1995.50: p. 3547-2556. Vaart, R. van der, Hafkamp, B., Koele, P.J., Querreveld, M., Jansen, A.E., Volkov, V.V., Lebedeva, V.I., Gryaznov, V.M., - Oxygen removal from water by two innovative membrane techniques, Conference on Ultra pure water Asia 2000. Mutto, T., Dean, D . - Development of Chlorine dioxide gas transport contactor (ERCO R101 Generator) for use in drinking water disinfection. Proceedings 12th Annual NAMS Meeting Lexington, KY, May 15-20 2001, p53. Majumdar, S., Bhaumik, D., Sirkar, K.K., Simes, G. - A pilot-scale demonstration of a membrane based absorption stripping process for removal and recovery of volatile organic compounds. Environ.Prog.,20(1) April 2001, p. 27-35. Celgard website: www. Celgard.com. Ho, W.S.W. and Poddar, T.K., - New membrane technology for removal and recovery of chromium from waste waters, Environ. Prog., 20(1) April 2001, p. 44-52.
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New Insights into Membrane Science and Technology: Polymeric and Biofunctional Membranes D. Bhattacharyya and D.A. Butterfield (Editors) © 2003 Elsevier Science B.V. All rights reserved.
Chapter 7
Membrane contactors: recent developments A. S. Kovvali + and K. K. Sirkar*
Department of Chemical Engineering New Jersey Institute of Technology, Newark, NJ 07102 +Present Address: Baker Petrolite, Sugar Land, TX, 77478; (Tel.) 281-276-5740; (Fax) 281-276-5492; email: [email protected] *Corresponding author: (Tel.) 973-596-8447; (Fax) 973-642-4854; email: [email protected] 1.
INTRODUCTION
Contacting two phases for efficient transfer of one or more solute species between the two phases is a primary objective of most conventional mass transfer operations. The contacting phase pairs can be gas-liquid, liquid-liquid, solid-liquid or solid-gas. The contacting area provided between the phases and the mass transfer coefficients of the solutes in each phase influence directly the transfer rates of solute species. Conventional phase contactors develop the contact area by forming of micro-phase domains (bubbles for gases, droplets for liquids and particles for solids) of one phase dispersed in the other phase. The interphase mass transfer rates are influenced strongly by the size and population of droplets, bubbles, particles etc. For superior rates of mass transfer to occur, the size of the droplets, bubbles etc. should be small while their number should be large. Another class of conventional contactors operates by forming thin films of one phase over which the other phase passes, accomplishing the species transfer. However, in each of the above mentioned cases, the dispersive contacting of phases necessitates the separation of phases after contacting. Further, scaling of such devices is problematic. Given such a context, membranes have found a niche role in phase contacting operations by providing the platform for nondispersive phase contacting. Simultaneously, they provide much larger value of volumetric mass transfer coefficient (kla, kga or ksa) over a much broader range of phase flow ratios without flooding, loading, channeling, weeping etc. Often such devices are called by a generic designation, membrane contactors; specific designations include membrane absorbers or membrane strippers or membrane extractors. Membrane contactors are gaining industrial recognition and commercial importance because of their unique features. Introduction to membrane 147
Membrane Contactors: Recent Developments - Sirkar
contactors and many of the unique advantages of membrane contactors have been identified in a number of book chapters and reviews [ 1-4]. New applications of membranes as phase contactors are being rapidly developed for a variety of applications. Membranes are being used or studied as contactors for: - physical transfer of one or more components from one phase to another phase in a nondispersive manner; - transfer of one or more components from one phase to the other phase where the transfer is enhanced by chemical reactions; - providing semipermeable barriers for one or more species while allowing contact between two phases for transfer of other species; and - providing complexing/reactive sites on a membrane surface for the transfer of one or more species from one phase, to be followed by elution/stripping. New and improved materials are being used in membrane contactors. As the range of applications expands, the demands on the membrane materials to withstand the conditions of the applications (temperature, pressure, chemical resistance) are different. New polymeric membranes are being used for membrane contactors. In some applications, the membrane contactor surfaces are modified to suit the particular hydrodynamic and chemical conditions imposed. Surface modification by forming a layer of another polymer is being attempted to provide stability to the phase interface and/or to provide chemical and fouling resistance rather than to provide any separation ability. Similarly, the ability of appropriate polymer chains grafted on membrane pore surfaces to provide large phase contact area for transfer of a species is also being explored [5]. As the demands on the membrane contactors grow, there is a need to optimize the contactor design to maximize its mass transfer capabilities. In this regard, new designs are being pursued, while some of the successful designs in other conventional areas are being adapted for new applications. This article attempts to review some of the key developments in membrane contactors in recent years; the emphasis will be on the new directions of membrane-based contacting technologies. The mechanistic aspects of the interface immobilization of two phases are discussed for their relevance to stable operation of membrane contactors. Recent advances in membrane materials and module configurations are discussed for their enhanced features. The applications of membranes as contactors can be broadly categorized as:
1) Liquid-gas contactors 2) Liquid-liquid contactors 3) Supercritical fluid-based contactors 4) Adsorption-based contactors 148
Membrane Contactors: Recent Developments- Sirkar
5) Phase contactors as reactors. Even though immobilized (supported) liquid membranes provide immobilized phase interfaces on each side of the porous/mesoporous/microporous membrane and may qualify as membrane contactors, the issues and the problems associated with their stability and performance are quite different from those relevant for membrane contactors. For this reason, they are not considered in this article. 2. MECHANISTIC ASPECTS OF THE INTERFACE IMMOBILIZATION IN MEMBRANE CONTACTORS Figure 1 presents a schematic representation of fluid-fluid interface in a membrane contactor. Unlike interfaces in conventional contacting equipment, the fluid-fluid interface and its location in membrane contactors is clearly defined. Fluid-Fluid Inerface ..q
Fluid 2
Fluid 1
I,
Polymer Support
Figure 1. Fluid-Fluid Interface in Membrane Contactors In most practical applications, the interface between two phases in a membrane contactor lies on one side of the membrane surface. It also implies that one phase usually occupies the porous region of the membrane. (However, this situation is quite different from "gas membranes" wherein the gas phase fills the pores of the membrane while the desorbing and absorbing solutions are in contact with the two surfaces of the membrane). Under these circumstances, the pore fluid (fluid 1) can not be displaced with the outside fluid (fluid 2) until the pressure difference between the external phase (fluid 2) and the phase in the pores (fluid 1) is greater than the breakthrough pressure (APb). The relationship is governed by Young-Laplace equation: ae~
=
2 y Cos O
(1)
r
where 7 is the interfacial tension between the two fluids, 0 is the contact angle measured from the pore wall to the tangent of the fluid-fluid interface and r is 149
Membrane Contactors: Recent Developments - Sirkar
the pore radius [ 1]. When 0 is zero, the substrate is completely wetted by the liquid. If a finite contact angle is formed (30 ° <0 < 89°), the substrate would be "partially wetting" and for 90 ° and above, the substrate is nonwetting. Moreover, the roughness of the surface plays a significant role in altering the contact angle and hence the breakthrough pressure [6]. A most important aspect in a successful operation of a membrane contactor is the ability of the phase interface to withstand as high a phase pressure difference between the two phases (AP) as possible without displacing the pore fluid. This is achieved mainly by a control of pore size and pore size distribution. As can be seen from equation 1, small and uniform pores in a membrane contactor ensure large and defined breakthrough pressures. The smaller the pore radius, the higher the breakthrough pressure. A narrow pore size distribution is helpful for phase interface immobilization in a given membrane contactor. For this reason, it appears that most membranes used in contactor applications are produced by techniques other than phase inversion and often by controlled stretching of melt extruded sheets/fibers or by thermally induced phase separation (TIPS) process. One aspect of the pore geometry to keep in mind is that the pore geometries in most membranes used for membrane contactors are considered as cylindrical. However, the actual pore geometry usually is non-cylindrical with slit like opening etc. Another important consideration in achieving a stable interface is the interfacial tension between the two fluids in contact. In liquid-liquid phase contacting, some applications involve very low interfacial tensions [7]. This has to be taken into account. For applications involving degassing of liquids, it is advantageous if the liquid does not wet the pores. In that scenario, ~ is equivalent to the surface tension of the liquid. The liquid will not enter the pores as long as the transmembrane pressure is less than the critical entry pressure. Young-Laplace equation govems this behavior as well [8]. An additional polymeric (usually dense and thin) layer can be deposited on the membrane surface; this layer will act as a physical barrier for the liquid to enter the pores while permeable species like volatile organic compounds (VOCs) permeate rapidly through such a coating [9]. Recently surface modification of the membrane has been attempted to address the wet-out problem. A specific example is a hydrogel layer to prevent wetting out of the membrane by oily components in fruit juice concentration by osmotic distillation [10,11]. As the thickness of this additional layer is usually small, the reduction in solute fluxes is expected to be minimum.
150
Membrane Contactors:
Recent Developments -
Sirkar
3. DEVELOPMENTS IN MATERIALS AND MODULE CONFIGURATIONS FOR MEMBRANE CONTACTORS The need for improved chemical resistance, appropriate wetting/nonwetting properties, good flow distribution in the contactor module, higher species transfer coefficients and reduced cost have led to a number of developments in substrate materials, membranes as well as module configurations. The wetting of a porous hydrophobic polymeric material by a liquid is determined by the difference in the surface tension of the liquid (~,) and the critical surface tension 0'c) of the polymer [ 12]. When ~/-~/ctends to zero, the membrane tends to get wetted. It is known that organic solutes in water reduce its surface tension. Further, organic solutes (volatile or nonvolatile) containing both polar and nonpolar functional groups may be adsorbed on the hydrophobic pore surfaces through the nonpolar functional groups and increase the value of ~,~ of the membrane material. Therefore a recent development has adopted the Goretex membranes of porous PTFE for gas absorption with aqueous solution of amines e.g. MEA, which are volatile since the value of ~ for polytetrafluoroethylene (PTFE) is quite low around 18 dynes/cm. Such a membrane does not get wetted with aqueous MEA solutions used to absorb CO2. Modules of such membranes are beginning to be used in large scale [13]. A second development utilizes conventional porous hydrophobic substrates (e.g. polypropylene, polysulfone etc.) but provides a thin coating of a high free volume fluoropolymer on one of its surfaces where the gas-liquid interface is located [14,15]. The nonporous coating on the membrane has sufficiently large free volume dimensions which provide very little resistance to species like H2S, CO2, N2, 02, CH4, H20, etc.; yet it essentially eliminates wetting of the porous substrate. The thin coating/skin of this new composite membrane provides an entirely new functional platform compared to the skin/coating of an asymmetric/composite membrane used in conventional membrane separation processes for separation of gas mixtures and liquid solutions. Since the notion of ozonation of aqueous streams was introduced by us using nonporous and microporous membranes with or without an intervening fluorocarbon phase [16-20], membrane-based ozonation is being practiced commercially. Contacting the highly reactive ozone with water in the presence of pollutants (e.g. semiconductor processing) creates a highly reactive environment. One needs chemical resistance, non-wetting membrane, higher species transfer coefficients amongst others. The new fluorinated polymer-based ozonator (INFUZOR) has all such features [21 ].
151
Membrane Contactors: Recent Developments - Sirkar
The choice of the polymer chosen as the membrane material, either for the substrate and/or for any coating, ultimately determines the long term stability of the membrane contactor for a given application. For most applications where the process streams are chemically compatible with polypropylene (PP) and where the liquid surface tensions are not too low, microporous PP hollow fibers appear to be the membrane of choice. Particular advantages with microporous polypropylene membrane contactors are their compactness (high surface area per unit volume) and their cost. In addition, new membrane materials are being developed to handle low surface tension liquids [22], harsh chemicals such as 03, NH3, HC1 [21] or to prevent wetting [23]. The applications for membrane contactors in medical applications essentially center around introduction of oxygen to blood and removal of CO2. The main additional feature required for these applications is the compatibility of the membrane material with the components of blood. Again, microporous PP is the membrane of choice [24]. Recent developments also focus on silicone-coated microporous fibers for preventing plasma leakage in long term use [25,26] The design of the membrane contactor also plays a significant role in the overall rates of species transfer and the membrane contactor size. For non-blood applications, LiquiCel Extraflow ® modules [22] appear to have become the standard configuration for membrane contactors. Currently transverse flow modules are available mainly with baffles. This configuration increases the mass transfer on the shell side significantly over earlier shell and tube membrane contactors [1]. Some of the new developments in membrane contactors using new membrane configurations [21,23] employ proprietary geometries to maximize the liquid side mass transfer coefficients. For medical applications, the primary issues in module design are low priming volume, low shear rates and low pressure drops [24]. For most applications involving membrane contactors, there are two steps involved: absorption of a solute from one phase into a second phase and desorption into a third phase to concentrate the solute. Examples of such applications are SO2 removal from gas streams [27], heavy metal removal from waste waters [28], isomer separation [29] using liquid membranes. Practically all these applications are pursued using two membrane contactors, one for absorption of the species and another for desorption (concentration). However, the two contactors can be integrated into one device with two sets of fibers. Such a configuration is called hollow fiber contained liquid membrane (HFCLM). An extensive review of this concept and applications was conducted by Sirkar [30]. One limitation of the concept is potential for maldistribution of the shell-side fluid, influencing the mass transfer rates of solute. In spite of the obvious advantages of compactness and economics, commercialization of this concept has not yet taken place. The reasons could be the difficulty in making a
152
Membrane Contactors: Recent Developments - Sirkar
device with two sets of fibers separated at the ports but interspersed intimately and uniformly in the device. 4.
CATEGORIES OF MEMBRANE CONTACTORS
Here we will identify different classes of contacting operations and briefly highlight important applications/developments. The different categories are listed in Table 1. Table 1. Categories of Membrane Contactors Category of Contactor Gas-Liquid Liquid-Liquid Supercritical Fluid-Liquid Liquid-Solid Contactors as Reactors All of the above categories of contacting operations except contactors as reactors are directed toward transferring of species from one phase to the other phase for the purpose of separation/purification whereas the last one utilizes such transport solely for the purpose of carrying out one or more chemical reactions. Generally both phases are mobile in the contactor except in the case of liquid-solid systems.
4.1
Liquid-Gas Contactors Gas-liquid contacting operations and devices are used to accomplish 1) degassing of liquids, 2) gassing of liquids, 3) absorption of gas/vapor, 4) gas/vapor stripping from absorbents, 5) membrane distillation, 6) supported gas membrane-based processes and 7) osmotic distillation. Such operations invariably use a porous hydrophobic membrane, most often in the hollow fiber form for large scale applications; the pores are filled with gas/vapor. In processes 1,2,3, and 4, the gas stream flows on one side of the membrane. The other side has a non-wetting liquid stream. In processes 5,6 and 7, both sides generally have a non-wetting liquid stream except in the processes of vacuum membrane distillation/air-gap membrane distillation. Tranverse-flow membrane contactors built out of hydrophobic microporous PP hollow fibers (Celgard ®) and possessing as much as 130 m E membrane surface area (based on 224640 fibers of O.D. 300 ~tm and I.D. 240 ~tm) in a module of dimensions 10" diameter and 28" long are routinely used to degas ultrapure water streams of dissolved oxygen; the liquid is in transverse flow on the shell side and the fiber bores are subjected to a combination of high vacuum and inert N2 purge [31]. The O2 level in the degassed water may be 153
Membrane Contactors: Recent Developments - Sirkar
reduced to less than 1 ppb. Another application of such devices/processes is to strip the volatile organic compounds (VOCs) from wastewater using air as the sweeping gas [32]. Although degassing of ultrapure water is one of the largest industrial applications of membrane contactors, the biomedical application of these types of hollow fiber membranes in blood oxygenators / artificial lungs of suitable design for removing CO2 from blood continues to be the largest degassing application. Concurrently, supplying 02 to blood via such membranes from a gas stream is the other major goal in the blood oxygenators. (A recent application of this technique is to use an intravenous membrane oxygenator [33] as an intracorporeal artificial lung as opposed to conventional extracorporeal devices). Gassing of streams in the same mode is a major thrust area of industrial membrane contactor applications. Current applications include carbonation of beverages [34], adding hydrogen to wastewater prior to its entering a membrane reactor [35] for removal of nitrates from drinking water via heterogeneous catalysis with catalysts embedded in a microporous polyetherimide membrane. Absorption of gases/vapors into an absorbent liquid in a membrane contactor for industrial gas separation/purification was one of the earliest areas of investigation [36]; the species to be absorbed were primarily CO2, SO2, NH3 amongst others. Absorption of CO2 in aqueous amines solutions containing monoethanolamine (MEA), diethanolamine (DEA) etc. in membrane contactors containing microporous hydrophobic PP fibers with/without a thin silicone coating encountered problems due to the wetting of the pores by the liquid within 7-15 days. Kvaerner in collaboration with W.L.Gore & Co have developed membrane contactors of porous polytetrafluoroethylene (PTFE) such that the membrane pores do not get wetted [ 13]. A large-scale CO2 absorption unit has been put up. These are particularly useful for CO2 removal in offshore applications due to drastic weight and height reductions. Klaassen and Jansen [37] have reported the installation of an ammonia absorption unit at a dyes intermediates production plant of Aliachem in Pardubice (Czech Republic). The ammonia-containing off gas stream is contacted with an aqueous solution which recovers NH3 from the gas stream as an aqueous ammoniacal solution (>20 wt%) to be reused in the dyes intermediate production process. The unit in operation since June 1999 has reduced NH3 emission to the environment by 99% and currently has a capacity to absorb 50 kg/hr. The volume and weight of the membrane absorbers were more than 10 times smaller than conventional absorbers. Additional applications being exploited include SO2 absorption [38]. The above applications involve liquids which are primarily aqueous and under ordinary circumstances many not wet the hydrophobic microporous membrane. Organic liquids (e.g. silicone oil) which are substantially nonvolatile and spontaneously wet the pores of such fibers have been
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successfully employed as an absorbent for the removal of VOCs from air streams in LiquiCel Extraflow ® membrane contactors under minipilot plant conditions using an actual paint booth exhaust stream [39]. This effort followed the laboratory scale studies of Poddar et al. [40] where they had used both porous membrane as well as a porous membrane with a nonporous silicone coating for absorption of VOCs in an oil. Stripping of absorbed gases from the liquid absorbent at a higher temperature with or without vacuum or an inert gas (e.g. N2) / vapor (steam) can also be implemented using membrane contactors. If a higher temperature is involved, special attention has to be paid to the materials of the membrane and the device. Efforts are ongoing with the microporous Goretex PTFE microporous membranes to strip the CO2 from the amine-absorption solution at a higher temperature [41]. Stripping of absorbed VOCs from the silicone oil absorbent was successfully implemented in a minipilot plant scale using vacuum and hydrophobic microporous polypropylene membranes having an ultrathin nonporous plasmapolymerized silicone membrane coating facing the silicone oil [39]. The coating was highly permeable to the absorbed VOCs but was impermeable to the silicone oil absorbent. Laboratory studies [9] carried out similar stripping successfully at a much higher temperature of 50-70 °C. The first successful laboratory studies of membrane distillation were by Gore [42]. Schneider et al. [43] studied the process of direct contact membrane distillation (DCMD) using larger units made from hydrophobic PP hollow fibers. Lawson and Lloyd [44] had demonstrated using very small fiat membranes that very high water vapor flux is possible in desalination of water by DCMD as much as 2-3 times that in reverse osmosis. However, this DCMD process is plagued by water vapor flux reduction and distillate contamination by brine with time due to wetting of the pores of the hydrophobic porous PP membranes normally employed. Further, high water vapor fluxes have to be achieved in a compact but high surface area contactor device. Vacuum membrane distillation process is not favored as such due to the need for a separate condenser. In DCMD, a hot saline solution produces water vapor which diffuses through the gas filled pores of a porous hydrophobic membrane and then condenses in the cold distillate water stream on the other side of the membrane. The water vapor pressure difference on the two sides of the membrane is the driving force through the gas in the pores, the so called "gas membrane" or the "supported gas membrane" (SGM). When there is a volatile species in an aqueous solution and an aqueous solution without the volatile species on the other side, then a porous hydrophobic membrane may be used to desorb the volatile species from the feed solution and get it absorbed in the receiving solution (for earlier references, especially the studies by Edward Cussler, see [2]). Chlorine dioxide (C102) has been stripped from an aqueous electrolytic 155
Membrane Contactors: Recent Developments - Sirkar
solution of an electrochemical reactor through a highly chemically resistant porous hydrophobic membrane stack and absorbed in the potable water to be disinfected and treated; large-scale plant operation has successfully taken place [45]. Osmotic Distillation is said to occur when the receiving solution on one side of the gas membrane is a highly concentrated saline solution having a vapor pressure of water which is much less than that of an aqueous feed solution on the other side of the membrane. Both solutions are at the same temperature just as in the C102-separation example above and unlike that in DCMD. Consequently, there is a partial pressure-based driving force for water transport from the feed to the receiving solution. Large-scale tests using two 19.2 m 2 LiquiCel ® modules and continuous production have been carried out for fruit juice concentration using such a technique [46]. Wetting of the pores due to hydrophobic components in the feed solution can be prevented by having a hydrogel layer on the surface of the hydrophobic membrane facing the feed fruit juice solution [ 11 ].
4.2
Liquid-Liquid Contactors Membrane-based liquid contactors are primarily employed to extract a nonionic or ionic solute selectively from one liquid phase into another immiscible liquid phase. The phases often encountered in an immiscible twophase system are aqueous and organic. (References to immiscible phase pairs which are organic-organic or biphasic aqueous are available in [1]). A large number of studies have been conducted exploring membrane solvent extraction of metals, organic acids, organic pollutants, antibiotics etc. from an aqueous solution into an organic extractant. Back extraction into an aqueous phase has also been conducted especially for metals, organic acids, antibiotics etc. A number of pilot plant studies were implemented. Almost all such studies employ Celgard hydrophobic microporous PP hollow fibers/Liquicel ® modules with the organic solvent phase in the pores and the aqueous phase outside at a higher pressure [47]. The Liquicel ® modules have the tube sheet made out of a lower Tg olefinic resin and have considerable solvent resistance. However, a number of solvents including chlorinated solvents, dimethyl formamide (DMF), dimethylsulfoxide (DMSO) etc. are not to be used. Back extraction studies of the solute into an aqueous phase also employ the same type of hydrophobic membranes/modules even though hydrophilic fibers with the aqueous phase in the pores is much more efficient from a mass transfer perspective [ 1]. A full-scale membrane-based solvent extraction plant has been operating since 1998 at Kosa, Vlissingen, Netherlands [37]. It employs Liquicel ® modules and an organic feedstock to extract an aromatic compound formed in a reactor
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and being discharged in an aqueous stream. The organic feed stock recycles the product back into the reactor. The waste water volume treated is 15 m3/hr. The waste water incineration facility was eliminated resulting in an annual saving of 5 million m 3 of natural gas. An additional example of large-scale hydrophobic microporous membrane-based solvent-extraction involves extraction of Cr 6+ from an aqueous solution into an organic extractant containing a secondary amine present in the pores of a Celgard fiber. The aqueous phase is maintained at a higher pressure than the organic extractant phase to immobilize the aqueous-organic interface in this plant operated by Commodore Separation Technologies, Kennesaw, GA [48,49]. However, there is a novelty in this process wherein the organic extractant phase has dispersed in it drops of the aqueous back extraction liquid which strips the Cr6÷ from the complex into a highly basic stripping solution. A similar technique is being used in large scale for removing zinc and iron from a passivating bath at Rogal, Enschede, The Netherlands [37]. A large enantiomer resolution plant at Tanabe Seiyuku, Osaka, Japan employs lipase enzyme-based resolution of the enantiomers present in an organic stream; the enzymes are reversibly immobilized in the pores of substrate of a hydrophobic asymmetric polyacrylonitrile (PAN)-based hollow fiber ultrafiltration membrane. The skin of this fiber is in the I.D.; an aqueous solution flows through the fiber bore. The aqueous-organic interface is immobilized on the fiber O.D. with the shell-side organic phase outside the pores maintained at a pressure higher than that of the aqueous phase. Total membrane area employed is 1440 m2. This technology is from Sepracor, Marlborough, MA. [50]; however the microporous membrane-based solvent extraction aspect of it was licensed from Hoechst Celanese Inc. (Current successor is Celgard Inc., Charlotte, NC). Additional techniques are being developed wherein liquid-liquid extraction takes place at one side of a porous membrane whereas a sweep gas/vacuum is maintained on the other side. It is called supported liquid membrane pervaporation (SLMPV) [51]. Highly stable technologies are currently being developed in our laboratory since conventional SLMs are generally considered unstable.
4.3
Supereritieal Fluid-based Contactors PoroCrit LLC, Berkeley, CA has started using Celgard fiber-based Liquicel ® contactors and other porous PP fiber-based devices to achieve species transfer from a liquid phase into supercritical CO2 or vice versa. They claim that "the process is conducted with the pressure on both sides of the membrane in the module being essentially the same" [52]. As the two phases flow in a countercurrent fashion in the PP hollow fiber module, they will undergo considerable pressure drops. So there will be considerable pressure difference 157
M e m b r a n e Contactors: Recent D e v e l o p m e n t s - Sirkar
between the two phases along the module length. Such an operation has been successfully used to extract flavors from wines and juices, recover aromas from vegetable and marine oils and nuts and fermentation broths and for industrial organic solvent recovery and recycle [53]. 4.4
Adsorption-based Contactors
When a gas or liquid phase containing a solute is contacted with a packed bed of adsorbents/ion exchange resin beads, adsorption or ion exchange-based transfer of solutes occur to the bed particles. There is significant~igh pressure drop in flow through packed beds, especially of fine particles. Further there is considerable resistance to diffusion of solute molecules (especially larger ones) through the porous resin beads; as a result, the utilization of the ligands etc. in the porous resin beads is low. Porous membranes used in the microfiltration separations and having pores in the range of 0.02-10 gm have been used as a substrate: ligands/biofunctional groups of appropriate types have been grafted onto the surfaces of pores in microfiltration membranes. The liquid phase containing solutes has been passed through the pores in the microfiltration mode of operation. The liquid convection through the membrane pores conveniently brings the solutes directly to the ligands on the pore surfaces eliminating the diffusional resistance encountered in the conventional resin beads. The process of adsorption/complexation takes place throughout the pore volume. Fractional ligand utilization is very high. The solute uptake process takes place rapidly. After some time, the feed flow is stopped and the ligands on the pore surfaces are regenerated by passing an eluent solution since they get saturated. Such devices are called membrane adsorbers but are in effect, liquidsolid contactors. Their application to protein separation is commercially practiced, albeit in a smaller scale. An earlier reference to the literature is available in Thommes and Kula [5]. Such a technique has now been adopted to heavy metal adsorption from waste waters with or without disposable membrane materials [54]. Disposable and cheap membrane materials (e.g. of cellulose) are useful especially if regeneration efficiency for the heavy metals is poor. 4.5
Contactors as Reactors
Whenever a reactor employs two immiscible phases which have to be contacted intimately, a membrane contactor may be employed. The two immiscible phases will flow/be on the two sides of the membrane and appropriate phase pressure differences have to be maintained for nondispersive operation. The reaction may take place in the bulk of one or both phases, at the interfaces on the pore mouths and/or in the fluid layers immediately adjacent to the phase interface [20]. Membrane-based ozonators developed by W.L.Gore and Pall Inc. are commercialized examples of such reactors: ozone-containing oxygen/air flows
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on one side of the membrane and the water to be treated flows on the other side. The materials of the device are chemically resistant perfluoropolymers. The membranes provide very little resistance to gas transport but do not usually allow liquid to come to the other side of the membrane. This area has an enormous potential for growth in the near future. The limitations are: selection of an appropriate example amenable to reaction in a membrane contactor; chemical resistance of the membrane material; controlling the thermal effects; flow maldistribution etc. Additional examples of membrane contactors employed as reactors are identified below: - Destruction of VOCs in a gas phase by their absorption in a fluorocarbon phase which is supplied with ozone from a gas phase via a separate membrane [55]. - Elimination of mercury from a gas phase by oxidative membrane absorption in an oxidizing aqueous liquid solution of H202/H2SO4 or K2Cr207 o r K 2 S 2 0 8 , K M n O 4 etc. [56]. The review paper by Sirkar et al. [57] provides additional examples especially in the production of chemicals/biochemicals. Amongst these examples are cases of hydrogenation of organic liquids via a porous ceramic membrane reactor which is essentially a membrane contactor. Porous ceramic membranes hold special promise especially due to their excellent thermal resistance. Membrane contactors have also been employed by Porocrit Co. with supercritical CO2 to sterilize juices and liquids by destroying microorganisms and inactivating enzymes [52]. 5.
MASS TRANSFER CORRELATIONS AND CONSIDERATIONS
The resistance to mass transfer of a species in a hollow fiber-based membrane contactor is the sum of three resistances: lumen side, pore and shell side. This can be represented as 1
=
1
1
1
~ + ~ + ~ (2) gtotat kl .... kpore kshett where Ktotal is the overall mass transfer coefficient, klumen, kpore, kshell are the local mass transfer coefficients on the lumen side, membrane pore, and shell side respectively. Depending on the application, one or more of these resistances may become negligible compared to the others. For example, for gas-liquid systems, gas phase resistance is usually negligible compared to liquid phase resistance. So, the side where the gas flows will not offer much resistance to mass transfer and almost all of the resistance is offered by the liquid side. For membranes wetted by the liquid phase, resistance offered by the liquid in the pore is constant and does not change with the hydrodynamics.
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In hollow fiber membrane contactors, lumen side mass transfer is usually well defined whereas the shell side mass transfer coefficient is a strong function of module geometry, fiber distribution and the resulting hydrodynamics. Recent efforts to understand the effect of shell side flows on the membrane contactor performance [31,58-60] highlight the complex nature of shell side hydrodynamics. In addition, these efforts may lead to a scaleup parameter for designing membrane contactors for different conditions. While membrane contactors are commercially available with different membrane areas [22] for handling a range of flow volumes, the blood oxygenation devices come in a limited number of sizes. Each device is presumably optimized for the required shear rates, pressure drops etc. 6.
FUTURE DIRECTIONS
Membrane contactors are finding increasing commercial acceptance in recent years. This trend is expected to continue with new applications attempted using membrane contactors. The successful use of membrane contactors will depend on one or more of its unique advantages (small footprint, flexible capacity, broad range of phase flow rates/ratios without associated problems, chemical resistance of materials etc.) which are not offered by conventional contacting equipment. The main class of applications where membrane contactors are used still appears to be for effective gas-liquid contacting in a nondispersive manner using aqueous streams. New applications in gas-liquid contacting and other areas are gaining ground. New materials and geometries are being used for making microporous substrates and membranes. The new materials are pursued mainly for their non-wetting properties and chemical resistance. Characterization of membrane contactors for their mass transfer is still empirical. A better understanding of the hydrodynamics (particularly shell side) and ways to quantify it in terms of mass transfer correlations is needed for effective scaleup of these contactors. Similarly, designs of membrane contactors for conventional applications are limited in number. As the range of applications grows, more designs are to be expected to cater to various aspects of contacting applications. 7.
ACKNOWLEDGEMENT
One of the authors (ASK) wishes to acknowledge many discussions with Dr. P.V. Shanbhag during the preparation of this manuscript.
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Sirkar, K.K., P.V. Shanbhag and A.S. Kovvali, Membrane in a Reactor: A Functional Perspective, Ind. Eng. Chem. Res., 38(10) (1999) 3715-3737.
[58]
Lemanski, J. and G.G. Lipscomb, Effect of Shell-side Flows on Hollow-Fiber Membrane Device Performance, AIChE J. 41 (1995) 2322-2326.
[59]
Lipnizki, F. and R.W. Field, Mass Transfer Performance for Hollow Fiber Modules with Shell-side Axial Feed Flow: Using an Engineering Approach to Develop a Framework, J. Membr. Sci. 193 (2001) 195-208.
[601
Schoner, P., P. Plucinski, W. Nitsch and U. Daiminger, Mass Transfer in the Shell Side of Cross Flow Hollow Fiber Modules, Chem. Eng. Sci., 53 (1998) 2319-2326.
164
New Insights into Membrane Science and Technology: Polymeric and Biofunctional Membranes D. Bhattacharyya and D.A. Butterfield (Editors) © 2003 Elsevier Science B.V. All rights reserved.
Chapter 8
Membrane Aromatic Recovery System (MARS) - A new process for recovering phenols and aromatic amines from aqueous streams .
•
F. C. Ferreira*, A. Llvlnt[ston
+q¢
, S. Han*, A. Boam +, S. Zhang
+
*Department of Chemical Engineering and Chemical Technology, Imperial College of Science, Technology and Medicine, Prince Consort Road, London SW7 2BY, Unity Kingdom; +Membrane Extraction Technology Ltd, Imperial College, London SW7 2BY, UK Corresponding author: Tel +44-20-75945582; fax +44-20-75945629, E-mail address: [email protected] 1.
INTRODUCTION
Phenolic compounds (chemicals such as phenol and its derivatives) are used in phenolic resins, polycarbonates, biocides and agrochemicals. Aromatic amines (chemicals such as aniline and its derivatives) are used in a wide range of consumer products, including polyurethane foam, dyes, rubber chemicals and pharmaceuticals. The factories that manufacture and/or use these types of chemicals often create aqueous waste streams containing significant concentrations (0.1-10wt%) of these chemicals. Both aromatic amines and phenolic compounds are toxic and many of them are also carcinogenic. Tightening legislation in many countries calls for dramatic reductions in emissions of these species. A variety of processes have been proposed for treatment of these aromatic amine or phenolic compound containing wastewaters. Off-site disposal (landfill, deepwell injection) and biodegradation result in the compounds, which have typical values in the range US$0.75-US$20 per kg, being lost. These compounds have high boiling points and low vapour pressures. Hence, processes that rely on liquid-gas phase transition, such as distillation and pervaporation, have high-energy requirements. The use of adsorbents, such as activated carbon [1-4] or resins [58], is usually expensive due to difficulties and complexity in the regeneration stage. Problems associated with the use of solvent extraction [9-13] arise with phase separation [14] and contamination of the wastewater with solvent [15] due to the intermediate polarity of the compounds, which require moderately water soluble solvents. 165
Membrane Aromatic Recovery System (MARS) - A New Process For Recovering Phenols And Aromatic Amines From Aqueous Streams Livingston
Emulsion liquid membranes [16-19] and surfactant liquid emulsion membranes [20-22]systems were used to extract aromatic acids and bases from wastewaters. In particular Li and co-workers [23] used an emulsion liquid membrane containing caustic as reactive agent to increase phenol mass transfer. The main drawback of the use of liquid membranes is its inherently instability due to the leakage of the internal phase to the wastewater and the easy swelling of the liquid membrane by the wastewater phase. To overcome this problem, porous membranes have been used to support liquid membranes [24]. To avoid breakthrough of the immobilized phase in the pores to the aqueous phase, Ho and et al. [25-29] report development of a supported polymeric liquid membrane, in which the pores of a microfiltration or ultrafiltration membrane are filled with polyamphiphilic polymeric (oligomeric) liquids. These membranes were successful used, for phenolic compounds extraction from wastewater to an alkaline stripping solution at lab scale. However the use of liquid membranes, including supported polymeric liquid membrane, always requires an additional step for membrane phase regeneration, which often requires difficult demulsifications and/or the use of additional chemical agents [16]. The use of a non porous solid dense membrane avoid this step, hence aromatic acids and bases extraction from wastewaters have been tried in systems such as pervaporation [30,31 ], and aqueous/membrane/organic perstraetion [32]. The use of an acid-base reaction to maintain the driving force for phenol [33] or aniline [34] dialysis through several membranes was demonstrated using, sodium hydroxide or sulphurie acid respectively. In these studies, polymethylsilane-polyearbonate eopolymer appears to be the more promising membrane, although the values for permeability appear lower than those for reported for PDMS (phenol 4.5x10 -11 mE.s-1 [35] o r 2 . 5 x 1 0 -11 mE.s-l[36] against lxl0 "11 mE.s-1 [33] and aniline 20.9x10 -11 mE.s-1 [37] against 2.5x10 "11 mE.s 1134]). In the work reported by Klein et al. [33,34], the molar concentration of phenolate or anilinium is kept much lower than the concentration of base or acid used as the stripping solution in order to ensure a high driving force, and the stated intention is to dispose of the extracted aromatics. In this present report, working under conditions where the molar concentration of the phenolate or anilinium is carefully maintained near to the concentration of acid or base, we are able to show that the recovery of relatively pure phases of phenol and aniline is possible. This chapter describes the Membrane Aromatic Recovery Systems (MARS), a stable membrane process able to recover valuable aromatic amines [37] and phenolic compounds [38] from aqueous streams and its successful application at pilot scale for aniline recovery. The aromatic compound is extracted from the k,wastewater to a stripping solution, through a nonporous membrane selectively permeable to aromatic in its non ionic form and
166
Membrane Aromatic Recovery System (MARS) - A New Process For Recovering Phenols And Aromatic Amines From Aqueous Streams Livingston
impermeable to charged species. The driving force is maintained using an acidbase reaction in the stripping solution, where the aromatic is accumulated in its ionic form until a high enough concentration to allowed aromatic recovery by neutralization and consequentially phase separation. The MARS process is shown schematically in Figure 1.
1.1
MARS Operating Principle Wastewater out stripped of dissolved Amines (phenols)
HCI (NaOH)
NaOH (HCl)
R-NH4 +
(R-O-)
,~!~i ~:~:, ,,i~,
R-NH4+ (R-O-)
Nonporous Membrane
~: ~
R-NH 3 recovered amines (ROH recovered phenols)
Saline aqueous phase Wastewater containing dissolved amines (phenols) Figure 1" Schematic diagram of MARS process showing operating principles. Diagram show configuration for aniline recovery- figures in brackets show phenol recovery configuration.
1.1.1 Extraction Stage The aromatic compound is continuously extracted from the wastewater through a nonporous membrane separating layer into a stripping solution, where pH is controlled. The stripping solution is acidic for aromatic amines, and alkaline for phenolic compounds. In the stripping solution, acid-base reaction takes place and the aromatic molecule is converted into an ionized form (anilinium chloride or sodium phenolate). The membrane separating layer is hydrophobic; therefore charged molecules (such as ionized aromatic rings) cannot cross back through the membrane into the wastewater. The driving force for mass transfer of the aromatic acid or base through the membrane from the wastewater is maintained because the aromatic reacts to form its ionised form, a chemically distinct species. Note that the pH control operates such that each mole of aniline or phenol diffusing across the membrane requires one mole of either acid or base for neutralisation in the stripping solution. 167
Membrane Aromatic Recovery System (MARS) - A New Process For Recovering Phenols And Aromatic Amines From Aqueous Streams Livingston
1.1.2 Recovery Stage The stripping solution is periodically collected and pH is adjusted in order to recover the nonionic form of the aromatic molecule by acid-base reaction (i.e. alkaline conditions for amines and acid conditions for phenolics). Anilinium and phenate (the ionized forms of the aromatic) have virtually infinite solubility in water. However the solubility of nonionic forms of aromatic amines and phenols are usually less than 5 wt%. Hence, when the aromatic ion (highly concentrated in the stripping solution) is neutralized to the nonionic form, the resulting concentration greatly exceeds the aqueous solubility limit, and the solution separates into two phases: an aromatic-rich phase and a saline aqueous underlayer. NaC1 is a by-product in both cases, but this salty underlayer can be simply recycled to the wastewater feed as shown in Figure 1. In many of the industrial wastewater samples studied by Membrane Extraction Technology Ltd, we have found that the salt created by the acid-base recovery mechanism is negligible compared to the concentration of salt already in the waste stream. For example, the MARS process typically produces around 0.5 g of salt per g of organic recovered. In a wastewater containing 10 g L -1 of aromatic, the process will remove the aromatic and add around 5 g L -~ of salt. Many chemical industry wastes contain in excess of 100 g L 1 inorganic salts which arise in reaction processes, so the additional salt is not usually an environmental issue. 2.
MATERIALS AND METHODS
Analytical techniques employed for concentration determinations are described elsewhere[37,38]. The membrane used was silicone rubber tube (70 controlle Pressure gauge(~ I
pH / probe
Acid or caustic for pH control stewater out ~
Wastewater container
Stripping Membrane tank with solution stirring and heating overflow
Membrane coil
Figure 2: Laboratory and pilot plant continuous extraction configuration, with wastewater inside membrane tubes.
168
Membrane Aromatic Recovery System (MARS) - A New Process For Recovering Phenols And Aromatic Amines From Aqueous Streams Livingston
wt% poly(dimethylsiloxane) or PDMS and 30 wt % fumed silica), supplied by Silex Ltd. The membrane tubes used in all the laboratory experiments and pilot trials were 3 mm internal diameter and 0.5 mm wall thickness, with the exception of results shown in Figure 4, where the membrane tube had 3mm internal diameter but 0.35 mm wall thickness. The continuous flow configuration used for extraction stage laboratory experiments and continuous extraction trials at Solutia, UK is shown in Figure 2. The batch extraction configuration used for trials at Solutia UK is shown in Figure 3. The membrane tank solution temperature was kept at 50°C in all laboratory experiments and pilot trials. pH controller pH probe
Stripping solution recirculation
Wastewater Membrane coil
Acid for pH control
Stripping solution Membrane tank with vessel, level rises stirring and heating as aniline transfers Wastewater outside and acid added membrane Figure 3" Pilot plant batch extraction configuration, with wastewater outside membrane tubes. 3.
LABORATORY RESULTS
3.1
Phenol and Aniline Extraction
In phenol experiments a 28 m length membrane tube was used, a synthetic wastewater containing 10g.L -1 phenol was fed to the process, and a 12.5wt% NaOH solution was used to control the pH of a stripping solution between pH 11-13. Figure 4 shows phenol inlet and outlet concentrations from the membrane tube and total phenol (phenol plus phenate) concentration in the stripping solution over time. It can be seen that total phenol concentration in the stripping solution increases from 0 g.L -1 to 200 g.L ~. Operation at steady state is shown in Figure 5, in which the total phenol concentration (phenol and phenate) in the stripping solution was 242 g.L-1, three times higher than the phenol solubility in water and twenty five times higher than phenol concentration in the wastewater. The measured phenol concentration corresponds well to the value
169
Membrane Aromatic Recovery System (MARS) - A New Process For Recovering Phenols And Aromatic Amines From Aqueous Streams Livingston
oo tO
~
30
i
25 =
i
8,--. =o
-
160~ Q-
i
O
-,~ "~15 o
200
~10
~
,,
;H-111iI RH- !
80 ~ .-
8e- ~-~40
5
E
~ o 120 8 ~
8 ~
u
(1)
e-
-
0
10
0
I
-o- inlet ~
20 30 Time (day)
40
outlet -,,- stripping solution I
Figure 4: Evolution of phenol concentrations in the inlet, outlet and stripping solution over time in the MARS process operated continuously at laboratory scale: non steady state (reprinted with permission from J.Membr.Sci[38], page 224).
calculated for steady state assuming 12.5 wt% NaOH is neutralised by phenol to pH 13 of 227 g L 1. Figure 4 shows outlet concentration increasing with increasing phenol concentration in the stripping solution. From calculations based on the pKa of phenol, at pH 11, 9% of the phenol in the stripping solution is present as nonionic phenol and 91% as phenate ion. Therefore, when the concentration in the stripping solution increases the nonionic phenol u} to
300
15-
t-
_J
Q.
t--
250 ==
=o310-
200
"5 .o_
8 ~
-150 pH=13
~
_
5-
-
_.¢ _c
100
50 0
|
0 o inlet
5
i
|
10
~
i
,
==
G) O
Q. 0.
~ o
"r.
~
""
I--
c
|
15 20 Time (day)
r~ outlet
~E
25
× stripping solution
I
I
Figure 5: Evolution of phenol concentrations in the inlet, outlet and stripping solution over time in the MARS process operated continuously at laboratory scale: steady state (reprinted with permission from J.Membr.Sci [38] page 228).
170
Membrane Aromatic Recovery System (MARS) - A New Process For Recovering Phenols And Aromatic Amines From Aqueous Streams Livingston
concentration becomes high enough to have a significant negative effect on the driving force. At pH 13 only 0.1% of the phenol is in the nonionic form, the driving force is restored, and outlet concentrations drop to lower values, as shown in Figure 4. In aniline experiments, performed at laboratory scale using a 10.5 m membrane tube length, a synthetic wastewater containing 5 g.L-1 aniline was fed to the process, and 10.5 wt% HC1 solution was used to control pH in the stripping solution at pH 1. - 250
108
-
-
(D "7 200 ~._ ~.-J 0
L
15o
~6
loo 2-
...._~ ~._..; 50
'
"~. 0 ~ ' =Q..
~
2.4 2.0
0
0 0
2O
40
60
80
100
I Time (day) o inlet "-" outlet ~stripping solution - - --Predicted value at steady sate based on HCI concentration fed Figure 6: Evolution of aniline concentrations in the inlet, outlet and stripping solution over time in the MARS process operated continuously at laboratory scale.
Figure 6 shows the aniline inlet and outlet concentrations in the membrane tube, and total aniline (anilinium and aniline) concentration in the stripping solution, increasing over time. At steady state total aniline concentration in the stripping solution was 218.5 g.L-1, six times higher than the aniline solubility in water and forty four times higher than aniline concentration in the wastewater. These values illustrate the potential of MARS as a concentrating process. The steady state aniline concentration also compares well to the value obtained from stoichiometric neutralisation of 10.5 wt% HC1, 212 g L -1.
After total concentration in the stripping solution reached a steady state, the pH set point was varied in order to study the effect of pH on mass transfer flux (after day 85). The higher the pH in the stripping solution, the higher the percentage of nonionized aniline, and therefore the lower the driving force, resulting in lower mass transfer rates. It can be seen in Figure 6 that an increase in pH of the stripping solution clearly corresponds to an increase in outlet aniline concentration.
171
Membrane Aromatic Recovery System (MARS) - A New Process For Recovering Phenols And Aromatic Amines From Aqueous Streams Livingston
In the laboratory experiments a short membrane tube was used in order to quantify the outlet concentrations and estimate the overall mass transfer coefficient. Temperature was used to enhance mass transfer. The overall mass transfer coefficient obtained at 50°C across the 0.5 mm thickness membrane, was 1.4x10-7 m.s-1 for phenol and 4.8x10 -7 m.s-1 for aniline. Each mole of aromatic extracted from the wastewater requires one mole of acid or base to be fed to the stripping solution for neutralisation, ie each mole of aniline extracted requires one mole of HC1 and each mole of phenol extract requires one mole of NaOH. These ratios were tracked during the experiments and were usually close (within + 10%) to 1.0.
3.2 Recovery Stage Aromatic recovery was performed in batch. The aromatic compound accumulated in the stripping solution in ionized form. The overflow of the stripping solution, enriched with aromatic in the ionized form, was fed to the recovery stage and neutralized in order to obtain the neutral form of the aromatic. Neutralization was performed by adding 37% HC1 solution to the sodium phenate solution and 50% NaOH solution to the anilinium chloride solution. The aqueous solubility of the aromatic in the neutral form is low (8.2 wt% for phenol and 3.4wt% for aniline), hence when neutralization is completed the solution separates into two layers: a top phase that is the final product, and a underlayer aqueous phase that contains resultant NaC1 and traces of aromatic.
80
n Phenol in the aqueous saline phase • .,.Phenol,in the orqanic phase /otat pnenot Ted-to recovery stage
,° t
60
50 -~
40
a.
20
30 10 0
1
Overflow period (day)
Figure 7: Mass balances for phenol batch recoveries from laboratory data (reprinted with permission from J.Membr.Sci. [38], page 225).
172
Membrane Aromatic Recovery System (MARS) - A New Process For Recovering Phenols And Aromatic Amines From Aqueous Streams Livingston
Mass balances for recovery of phenol are shown in Figure 7. The organic phases were composed of aromatic compound and water. With phenol, the organic phase consists of 86.5% phenol and 13.5% water and, with aniline, 96.5% aniline and 3.5% water. No other organic was detected by GC analysis in the laboratory experiments. Recovery efficiency is defined as the ratio of mass of aromatic in the organic phase over the mass initially present in the stripping solution. Recovery efficiencies of over 92% were achieved for both systems. The balance of aniline or phenol is present dissolved in the aqueous layer resulting from the recovery process. The NaC1 in this saline aqueous underlayer is a by-product, which is responsible for a salting out effect that limits the solubility of the aniline or phenol in the aqueous layer. Hence, the higher the NaC1 concentration in the aqueous layer, the lower the aromatic concentration in the saline underlayer and the lower the water concentrations in the organic phase. Therefore the higher the NaC1 concentration, the higher the aromatic concentration in the final product and the higher the recovery efficiencies become. Figure 8 shows the salting out effect, with phenol aqueous solubility decreasing from 83 to 15.7 g.L1 as NaC1 concentration increases from 0 to 200 g.L 1 and phenol concentration in the organic layer increasing from 70.5 to 87.5 wt% for the same range of NaC1 concentration. Clearly there is a benefit to operating by adding as high a concentration of acid or base to the stripping solution as possible. However, this is balanced by the overall concentration of phenol or aniline rising and increasing the concentration of unionised aromatic in the stripping solution, leading to lower mass transfer fluxes. .~"~ 90 80 ; 70 c
k \
e Solubility o Phenol in organic _ phase
1 100 | | ~"
60
90
50
Q..
.~
~. 4o
"6
80 o
30
_c-
~= 20 10 0
0 ¢-
•c--
,
,
,
100
200
300
70 n 400
NaCI in water (g.L -~)
Figure 8: Phenol concentration in the aqueous phase, water and in the organic-rich phase as a function of NaC1 concentrations in the aqueous phase (reprinted with permission from J.Membr.Sci. [38], page 227). 173
Membrane Aromatic Recovery System (MARS) - A New Process For Recovering Phenols And Aromatic Amines From Aqueous Streams Livingston
3.3
Other Aromatics
In order to illustrate the applicability of MARS to recover compounds other than aniline and phenol, the overall mass transfer coefficients of a range of compounds were measured and results are shown in Table 1. Three independent non-coiled (randomly spaced) membrane tubes were submerged in a single stripping solution vessel, pH was kept below 1 using HCI solution for aromatic bases and over 14 using NaOH for aromatic acids. Under these conditions the unionised aromatic compound concentration in the stripping solution was assumed to be zero. In each run, three aromatic compounds were tested. The overall mass transfer coefficient for aniline and phenol in Table 1 are around twice (respectively 1.7 and 2.2 times higher than) the average value obtained in the laboratory MARS extraction. This is probably since in the mass transfer test the membrane was loosely arranged whereas the membrane in the MARS process was tightly rolled around a support, reducing the effective membrane-stripping solution interfacial area. Dimethylamine,4-nitrophenol,and 2,4,6 tris(dimethylaminomethyl)phenol exhibit very low overall mass transfer coefficients and their extraction from wastewaters will be difficult using silicone rubber tubes as membranes. Mass transfer rates for phenol, 4-nitroaniline, hydroquinone and pyridine have Table 1" Overall mass transfers for different compounds at 50°C across a 0.5 mm wall thickness silicone robber membrane Compounds
Kov x 107 (m.s-1)
Compounds
Kov x 107 (m.s"1)
Aniline 4-chloroaniline 2,4-chloroaniline 4-nitroaniline 4-fluoroaniline 2,4-fluoroaniline triethylamine dimethylamine benzyldimethylamin dicyclohexylamine
8.2 11.6 6.36 4.34 10.27 9.33 20 0.72 18 16.5
phenol 4-chlorophenol 2,4-dichlorophenol 4-nitrophenol 4-cresol hydroquinone Pyridine 2,4,6- tris (dimethy laminomethyl) phenol
3.1 9.3 14.7 0.48 7.04 2.37 4.4 0.64
intermediate values, and the extraction of these three compounds from wastewaters by MARS technology is possible using silicone rubber as membrane, but will use relatively large membrane areas. All other compounds tested have higher overall mass transfer coefficients, and so they can be
174
Membrane Aromatic Recovery System (MARS) - A New Process For Recovering Phenols And Aromatic Amines From Aqueous Streams Livingston
relatively easily removed from wastewaters by MARS technology using silicone rubber tubes as membranes. 4. PILOT PLANT RESULTS FOR ANILINE The MARS process has been scaled-up for recovery of aniline at Solutia, UK by Membrane Extraction Technology Ltd (UK). Further pilot trials are currently underway with phenol. Production of 4-nitrodiphenylamine (4NDPA) at Solutia UK results in a wastewater containing an aniline concentration that varies from 4.1 to 9.5 g.L -1, with an average value of 6 g.L 1. A pilot scale MARS unit was operated on site and recovered a total of 120 L of aniline from 29,500 L of wastewater over a period of some two months intermittent operation. 4.1
Continuous Extraction
Initially the pilot plant was configured for continuous operation, as for laboratory scale work (Figure 2), in which the wastewater flowed inside the membrane tube and aniline was accumulated out side of the tube, in an acidic stripping solution. At lab scale a membrane area of 0.099 m 2 (a single 10.5 m length tube) was used at a flow rate of 5.5 L.day -1, immersed in a membrane tank holding 1.5 L of stripping solution with pH controlled at 1. The pilot plant used a total area of 45 m" (48 tubes with 100m length), immersed in a 1000 L membrane tank and a pH in the stripping solution of 1.5. Flowrate was between 200-500 L.day 1. The temperature of the stripping solution was kept at 50°C and an 11% HC1 solution was fed to the stripping solution to control pH. Quantities of the stripping solution (80 L in the pilot plant against 0.25L in the lab scale) were periodically sent to batch recovery, which was achieved by addition of NaOH (33 wt% in the pilot plant, against 50 wt% in the lab scale). 12 11 10
60 - * - - i n l e t - - e - o u t l e t -e-stripping solution I
50
a) .-.. 9
=-~'
~
= o
6~
40 "~ ~, 88
8
"-"
7
o
15
30
4
20
2
10 ~-';
o ._ -~. m
~
1 0
.
.
.
.
,
5
.
.
.
.
.
.
10
.
.
.
.
.
15
.
.
.
.
.
.
.
.
,
20 25 Time (days)
.
.
.
.
.
.
30
.
.
.
.
.
.
35
.
.
i
0
40
Figure 9: Evolution of aniline concentrations in the inlet, outlet and stripping solution over time in the MARS process operated continuously at pilot plant scale.
175
Membrane Aromatic Recovery System (MARS) - A New Process For Recovering Phenols And Aromatic Amines From Aqueous Streams Livingston
Figure 9 shows aniline inlet and outlet concentrations, and total aniline concentration in the stripping solution, over time during continuous pilot plant operation. The average remo~cal achieved was 95% over the whole operating period, which corresponds to an overall mass transfer coefficient of 2.7 x 10-7 m.s 1, somewhat lower then the 4.8x10 -7 m.s -1value obtained at lab scale. 4.2
Interruption of Continuous Extraction The continuous operating period ended after day 36 when the backpressure of the wastewater being pumped down the membrane modules rose suddenly to over 0.5 bar. This caused a relief valve in the wastewater line to open and flow through the membranes ceased. Examination revealed that the membranes, and pipework downstream from the membranes, had blocked. The membrane tubes were removed from the membrane tank and some of them were opened up to reveal crystals of a black material, which were obviously blocking the flow. This black material dissolved readily in dichloromethane. When injected onto the GC, it was identified as 84% (by peak area) 4NDPA. The blockage corresponded to a dark-coloured organic phase being present in the wastewater due to insufficient upstream phase separation. This organic phase is probably organic product phase from the 4NDPA reaction. This mixture of aniline, 4NDPA and other compounds was initially liquid (aniline acts as a solvent for 4NDPA). In passing through the membranes, the aniline was stripped leaving the 4NDPA originally present in the liquid mixture as crystals in the membrane tubes. Filtration methods to avoid entry of this organic phase in the membrane tube were attempted, however frequent cleaning and filter changes are required and the filters were unable to prevent the passage of liquid solvents, and so were not sufficiently reliable or effective. 4.3
Batch Extraction To overcome these practical challenges, the pilot plant was reconfigured to extract aniline from wastewater batches, with the acidic stripping solution flowing inside the membrane tube and the wastewater in the membrane tank as shown in Figure 3. 770 L of wastewater was treated in each run. The same membrane area was used for batch and continuous extraction, and the membrane tank was kept at the same temperature (50°C). pH in the stripping solution was controlled at 1.5 using a 11% HC1 solution. A re-circulating flow rate of 80 L.h -1 was used in the stripping solution. In this configuration the wastewater, with the tar or organic phases when they are present, were pumped directly to the membrane tank without filtration. Membrane blockage was avoided since the solution inside membrane tube is the stripping solution in which no blockage material was present. Figure 10 shows aniline concentration initial and final
176
Membrane Aromatic Recovery System (MARS) - A New Process For Recovering Phenols And Aromatic Amines From Aqueous Streams Livingston
100
~'6o
- BO
°ii iI
r~
4"~
!
I
20
0
1 2 3
4
5 6
7 8
9 10 11 12 13 14 15 16 17 18 19 20 21
/ A n i l i n e inlet D Aniline outlet--*-Aniline removal
Batch Number
Figure 10" Batch operation pilot plant data: aniline initial and final concentration in each batch, aniline removal in each batch. concentration for each of the batch runs and the respective removal efficiency. Batch extraction was on average 22 hours in duration. An average of 90% removal over all the 21 batches was achieved. Aniline removal can easily be increased by increasing the membrane area/ membrane tank volume ratio (values of 100 m2m3 are easily obtainable). The stripping solution aniline concentration increases from 80g.L -1 to 160 g.L -1. The theoretical steady state of 240 g.L -1 which corresponds to 11% HC1 solution used was not reached in the time span of the trial. An average value for overall mass transfer coefficient of 4.2 xl0 -7 m.s 1 was calculated during the pilot plant batch extractions. This value is comparable to the average value obtained in laboratory scale continuous extraction (4.8x107 m.s-1), and higher than 2.7 x 10-7 m.s 1, the value obtained during continuous operation of the pilot plant. The mass transfer rate in the continuous pilot plant most probably decreased due to partial blockage of some tubes by tarry solids.
4.4
Recovery Stage
Aniline recovery from the acidic stripping solution enriched in anilinium chloride was achieved by neutralization using a 33% NaOH solution. A total of 128 L of aniline rich phase was recovered with a composition of 95 wt% aniline, 2.4 wt% Toluidine and 2.6 wt% water. The average aniline concentration in the salt under layer is 4 g.L -1. The purity of aniline obtained was high enough to allow it to be re-used in 4NDPA production.
177
Membrane Aromatic Recovery System (MARS) - A New Process For Recovering Phenols And Aromatic Amines From Aqueous Streams Livingston
5. PROCESS ECONOMICS Aniline price at time of writing is about $0.75 Kg -~, and hence it is not one of the more valuable aromatics in the market. Nevertheless due its Item
Comment
MARS Plant
Capital Charge Factor = 0.3 $45,000
Annual Cost (Benefit)
Membrane Replacement 2 year lifetime
$7,500
Acid (33% HC1)
17 ty-l@ $105 t-1
$1,785
Base (50 %)
12 t y -1 @ $210 t -1
$2,520
Power ( 3 kW)
8000 h y-~ @ $0.075
$1,800
Steam
82 t y-1 @ $15 t-1
$1,230
Labour
10% of one staff
$11,250
TOTAL
BENEFIT
($41,415)
Table 2: Case S t u d y - Process economics of MARS process using fluoroaniline as an example.
environmental impact, aniline removal is necessary and MARS provides a considerable economic advantage to alternative processes, although details cannot be provided here for commercial reasons. However, an example of MARS to recover a more valuable chemical is given in the Case Study contained in Table 2. The commercial value of fluoroaniline is ten times the aniline value (around $7.5 kg-1). In the example given the treatment of 10 m3 d-1 wastewater with a 5 g L 1 fluoroaniline concentration is considered. The MARS process in this case is able to deliver a net benefit through the value of the recovered fluoroaniline. 6. CONCLUSIONS MARS is a novel process coupling detoxification and recovery. It is capable of achieving high recovery efficiencies, and producing a relatively pure stream of recovered organics. At the laboratory scale, MARS has proven to be a successful process for removing and recovering aniline and phenol, with a recovered purities of 95% and 86.5% respectively, pH is an important parameter which controls the driving force across the membrane and consequently the removal efficiency of the aromatic from the wastewater. Anilinium and phenate salts were accumulated in the stripping solution to high concentrations, allowing further recovery by neutralization. The strip feed solution (HC1 or NaOH) concentration is a key parameter that controls the aromatic concentration at steady state in the stripping
178
Membrane Aromatic Recovery System (MARS) - A New Process For Recovering Phenols And Aromatic Amines From Aqueous Streams Livingston
solution. The salting out effect, resulting from high concentrations of NaC1 in the aqueous phase after neutralization, has a positive effect on the phase separation, decreasing the water concentration in the organic phase and the aromatic concentration in the aqueous phase. At pilot plant scale MARS was proven able to recover aniline from an industrial wastewater in a good purity. MARS was shown to be easily scaled up based on membrane area. Using different configurations, MARS was adapted to deal with a common problem in membrane technology, that is membrane blockage by tarry solids or organics precipitation. The MARS process can utilise very simple nonporous rubber tubes as membranes. Mass transfer is relatively low, but these membranes are relatively cheap. This is a key area where we expect to improve the process over the next months. MARS has low energy requirements because it exploits the acid-base functionality of aromatic acids and bases to produce a driving force based on the chemical energy contained in NaOH or HC1. The process can be carried out at conditions of pressure and temperature which are near ambient throughout all items of equipment. MARS does not rely on volatility of organics (or any phase transition), so can recover organics which membrane technologies such as pervaporation cannot reach. Finally, MARS has been shown to be promising for industrial application to recovery of a larger range of aromatics than phenol and aniline. Examples include amines, phenolics and pyridines. REFERENCES [ 1] [2] [3] [4] [5] [6] [7] [8] [9] [10]
ChiangP.C., Chang E.E., Wu J.S.: Comparison of chemical and thermal regeneration of aromatic compounds on exhausted activated carbon. Wat. Sci. Tech. 35-7 (1997) 279. OrshanskyF., Narkis N.: Characteristics of organic removal by PACT simultaneous adsorption and biodegradation. Wat. Res. 31-3 (1997) 391. Cha T., Glasgow: Sorptive reclamation of phenol from coal conversion wastewater. Ind. Eng. Chem. Res 22 (1983)198. Pahari P.K., Sharma M.M.: Recovery of heterocyclic amines from dilute aqueous waste streams. Ind. Eng. Chem. Res.30 (1991) 1980. Fox C.R. Plants uses prove phenol recovery with resins. Hydrocarbonprocessing 11 (1978) 269. Fox C.R. Remove and recover phenol. Hydrocarbon processing 7 (1975) 109. Fox C.R.: Removing toxic organics from waste water. Chem. Eng. Process. 8 (1979) 70. FergusonG.U.: Phenol recovery using polymeric adsorbent resins. Chem. Ind.8 (1982) 567. TeramotoM., Takihana H., Shibutani M.Yuasa T., et al: Extraction of amine by W/O/W emulsion system. J. Chem. Eng. Jap. 14-2 (1981) 122. Jagirdar G.C., Sharma M.M.: Separation of close boiling substituted Anilines: Gasliquid vs Conventional liquid-liquid dissociation extraction. J. Separ. Proc. Technol.24(1981)7.
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[11]
[12] [13]
[14] [15]
[16] [17]
[18] [19]
[201 [21]
[22] [23]
[24] [25] [26] [27] [28] [29]
[301 [31]
Wadekar V.V., Sharma M.M.: Separation of close boiling organic acids/bases by dissociation extraction: Binary and Ternary Systems: Substituted anilines; Binary system with thermally regenerative extractant: Chlorophenols. J. Sep Proc. Tech.2-2 (1981)28. Medir M., Arriola A., Mackay D., et al: Phenol recovery from water effluents with mixed solvents. J. Chem. Eng. Data 30 (1985) 157. Jagirdar G.C., Sharma M.M.: Recovery and separation of mixtures of organic acids from dilute aqueous solutions. J. Separ. Proc. Technol. 1-2, (1980) 40. Netke S.A., Pangarkar V.G.: Extraction of naphthenic acid kerosene using porous and nonporous polymeric membranes. Sep. Sci. Technol. 31 (1996) 63. Krishnakumar V.K., Sharma M.M.: A novel method of recovering phenolic substances from aqueous alkaline waste streams. Ind. Eng. Chem. Process Des. Dev. 23(1984)410. Devulapalli R., Jones F: Separation of aniline from aqueous solutions using emulsion liquid membranes. J. Hazard. Mater. B.70 (1999) 157. Boyadzhiev L., Besenshek E., Lazarova Z.: Removal of phenol from waste water by double emulsion membranes and creeping film pertraction. J. Memb. Sci.21 (1984) 137. Thien M.P., Hatton T.A.: Liquid emulsion membranes and their applications in biochemical processing. (2001) 819. Gadekar P.T., Mukkolath A.V., Tiwari K.K.: Recovery of nitrophenols from aqueous solution by a liquid emulsion membrane system. Sep. Sci. Technol. 27 (1992) 427. Kataoka T., Nishiki T., Kimura S.: Phenol permeation through liquid surfactant membrane-Permeation model and effective diffusivity. J. Memb. Sci. 41 (1989) 197. Kakoi T., Goto M., Natsukawa S., et al: Recovery of phenols using liquid surfactant membranes prepared with newly synthesized surfactants. Sep. Sci. Technol.31-1 (1996) 107. Ulbrich M., Marr R., Draxler J.: Selective separation of organic solutes by aqueous liquid surfactant membranes. J. Memb. Sci.59 (1991) 189. Terry R.E., Li N.N., HO W.S.: Extraction of phenolic compounds and organic acids by liquid membranes. J. Memb. Sci. 10 (1982) 305. Kiani A., Bhave R.R, Sirkar K.K.: Solvent extraction with immobilized interfaces in a microporous hydrophobic membrane. J. Memb. Sci.;20 (1984) 125. Ho S.V.: A supported Polymeric Liquid Membrane Process for Removal of Carboxylic Acids from a Waste Stream. Environ. Prog. 18-4 (1999) 273. Ho S.V. Extracting organic compounds from aqueous solutions. (1996) US. Patent 5512180. Ho S.V. Solid poly-amphiphilic polymer having use in a separation process. (1996) US. Patent 5,552,053. Ho S.V. Supported liquid membrane and separation process employing some. (1996) US. Patent 5507949. Ho S.V., Sheridan P.W., Krupetsky E.: Supported polymeric liquid membranes for removing organics from aqueous solutions I. Transport characteeristics of polyglycol liquid membranes. J. Memb. Sci. 112 (1996) 13. Meckl K., Lichtenthaler R.N.: Hybrid process using pervaporation for the removal of organics from process and waste water. J. Memb. Sci. 113 (1996) 81. Lipnizki F., Hausmanns S., Ten P.K., et al: Organophilic pervaporation: prospects and performance. Chem. Eng. J. 73 (1999) 113.
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[32] [33] [34]
[351 [36]
[37]
[38]
Ray S.K., Sawant S.B., Joshi J.B., et al: Perstraction pf phenolic compounds from aqueous solution using a nonporous membrane. Sep. Sci. Technol. 32-16 (1997) 2669. Klein E., Smith J.K., Weaver R.E.C., et al: Solute Separation from Water by dialysis II. The separation of Phenol by Downstream Conjugation. Separ. Sci. 8-5 (1973) 592 Klein E., Smith J.K., Wendt R.P., et al: Solute Separation from Water by dialysis I. The separation of Aniline. Separ. Sci. 7-3 (1972) 285. Brookes P.R, Livingston A.G.: Aqueous-aqueous extraction of organic pollutants through tubular silicone rubber membranes. J. Memb. Sci. 104 (1995) 119. Bennett M., Brisdon B.J., England R., et al: Performance of PDMS and organofunctionalised PDMS membranes for the pervaporative recovery of organics frm aqueous streams. J. Memb. Sci.137 (1997) 63. Castelo-Ferreira F., Han S., Livingston A.G.: Recovery of Aniline from Aqueous Solution using the Membrane Aromatic Recovery System (MARS). Ind. Eng. Chem. Res. (2002) In Press. Han S, Castelo-Ferreira F., Livingston A.G.: Membrane aromatic recovery system (MARS) -a new membrane process for the recovery of phenols from wastewaters. J. Memb. Sci. 188-2 (2001) 219.
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Functional Membranes and Materials for Biocatalysis, Separation, and Analysis
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Biofunctional Membranes and Biomedical Devices
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New Insights into Membrane Science and Technology: Polymeric and Biofunctional Membranes D. Bhattacharyya and D.A. Butterfield (Editors) © 2003 Elsevier Science B.V. All rights reserved.
Chapter 9
Membrane bioreactors for biotechnology and medical applications L. Giorno, L. De Bartolo, E. Drioli* Research Institute on Membranes and Modelling of Chemical Reactors, IRMERC-CNR, c/o University of Calabria, via P. Bucci cubo 17/C, 87030 Rende (CS); *Corresponding author, Tel: +39 0984 492039; Fax. +39 0984 402103 e-mail: [email protected] ABSTRACT In this chapter we will discuss the properties and applications of bioartificial hybrid systems used in biotechnological and medical applications. In the first part we will focused on membrane bioreactors using biocatalysts such as enzymes, microrganisms and cells for the production of biological molecules or for separation of pharmaceutical products. In the second part we will discuss on membrane bioreactors using isolated mammalian cells (i.e., liver cells, pancreatic cells) as bioartificial organs in temporary or continuous substitution of injured organ. The properties of membranes to be used in these devices will be reported including mass transport, morphological and physico-chemical properties that influence their performance. The discussion of the biotechnological part will consider the effect of immobilization on functional stability and activity of biocatalysts, including membrane material and morphology, physico-chemical properties of reaction environment. The successful examples of bioreactors running at large scale, of which the authors are aware, will be also presented. In the medical application part, the development of membrane bioartificial organs, e.g. bioartificial pancreas and bioartificial liver, as well as the properties of membranes for these systems will be discussed. Particular attention will be given to the recent achievements in cytocompatibility and biocompatibility of membranes in bioartificial organs. 1.
INTRODUCTION
The combination of artificial membranes with macromolecules of biological origin allow to obtain hybrid systems capable to imitate what nature
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has fine tune through aeons of evolution. In fact, these systems, based on their selective transport and catalytic properties, are able to carry out chemical reactions, selective separation, waste removal, etc. The use of membranes and biological tools for improving traditional production systems is a possible approach for maintaining a sustainable growth. Typical examples include novel pharmaceutical products with well-defined enantiomeric composition; new foods with improved nutritive properties; medical devices for the care of crucial diseases; and the treatment of wastes. The performance of the hybrid systems are governed by the properties of specific biomolecules and membranes and in particular by their reciprocal interactions [ 1]. When the biomolecule is an enzyme immobilized in a membrane, the overall performance of the biocatalytic membrane reactor is due to the preserved functional conformation of the enzyme as well as to the transport of reagent and product, through the membrane, to and from the enzyme, respectively. When mammalian cells are to be adsorbed on membrane to prepare artificial organs, the membrane properties are of crucial importance to keep them viable. The biomolecule can either be loaded to the membrane or just compartmentalized by the membrane. Reactors with segregated cells are suitable for therapeutic applications; for example, the bioartificial pancreas, the bioartificial liver, and the extracorporeal detoxification device [2]. 2.
M E M B R A N E BIOREACTORS IN B I O T E C H N O L O G Y
Biocatalytic membrane reactors can combine selective mass transport with chemical reactions and the selective removal of products from the reaction site increases the conversion of product-inhibited or thermodynamically unfavorable reactions. Many enzymes and micro-organisms have been used in membrane reactors to catalyze bioconversions for various application [4-24]. The possibility of integrating biotransformation into productive reaction cycles makes reliable the use of biocatalysts at large scale. The main parameter that influenced the development of biocatalytic processes at productive level is the stability of the biocatalyst. Immobilization has proved to increase stability of biocatalysts, although it may cause changes in the catalytic activity and enantioselectivity [25-27]. This is a common observation, nevertheless, it cannot be considered a general rule of the inverse relationship between stability and activity and enantioselectivity of immobilized enzymes. Most probably, these effects are related to the interactions between the chemical groups of enzyme and membrane: due to these interactions the molecule becomes more rigid, and therefore more stable, but the loss in flexibility damages the specificity and activity. 188
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Recently, studies on the structure of enzymes present in extremophiles (thermo-stable) and mesophiles (thermo-labile) micro-organisms have been reported [28-31 ]. Referring to the activity, it is observed that the enzyme that are resistant at high temperature are much less active (especially at low temperature). For immobilized enzyme, the challenge is to obtain the proper interactions to guarantee a balanced efficiency of the various properties (stability, enantioselectivity and activity). The efficiency of the overall immobilized system, depend on the biochemical (catalytic activity; half-life time; productivity; reaction kinetics; catalyst inhibition effects; concentration, viscosity, solubility and purity of substrate and product; sensitivity to pH and temperature; immobilization stability, etc.) mechanical (membrane configuration, structure and strength; size, morphology and pore size distribution; swelling and compaction effects; shear stress, etc.) and hydrodynamics parameters (flow regime; pressure drop; transmembrane pressure; flow velocity; residence time; easy to clean, etc). A comparative analysis of the performance of enzymes free and immobilized on membrane using various substrates is reported in Table 1. Hydrolases/esterases (such as lipases) are among the most studied biocatalysts [34-50]. Lipases (triacylglycerol-hydrolases, E.C. 3.1.1.3) act efficiently on water-insoluble substrates by binding to the water-organic interface. The natural organic interface consists of triglycerides. This binding places the lipase close to the substrate, and in particular increases the catalytic efficiency by interfacial activation. Due to their natural function at the oil-water interface, lipase exhibit high stability in contact with organic solvents, and beside triglycerides they are also active with a variety of non-natural substrates, showing generally a high stereo- and/or regioselectivity [41, 53]. Some author reported that the immobilized lipases used in organic solvent to catalyze esterification reactions, show higher activity and stability than lipase free in suspension [54, 55]. 2.1
Influence of operating conditions
2.1.1. Effect of immobilization
Lipases have on their surface hydrophobic and hydrophilic patches, therefore they can be immobilized by ionic interactions on ionic hydrophilic membrane support, or by van der Waals forces on hydrophobic membrane. Hydrophobic materials have shown to increase the performance of lipase [63, 64]; nevertheless, the support has to be chosen in agreement with the properties of the overall reaction system. For example, when using organic solvents, the resistance of the membrane to the solvent has to be taken into the account. In this case, hydrophilic membranes (such as polyamide,
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polyacrilonitrile) have shown to have higher stability than hydrophobie ones (e.g., polypropylene). Table 1" A comparative analysis of the performance of enzyme immobilized in various
membrane Enzyme (substrate)
Lipase (Olive oil)
Type of immobilization
Type of reactor
(Enzyme Free) a
STR E-EMR
Observed reaction rate (retool.1 "l h q)
Enzyme stability
ef.
-
6.95
Low
2
1.6
6.80
Medium
4.50
High
Axial Flow rate (ml-min "t)
3, 34
Entrapment by cross-flow filtration
TSP-EMR (organic) 1.6 (aqueous) 160 Cross-linkingby TSP-EMR (organic) 1.3 glutharaldehyde (aqueous) 167 Lipase (Cyano ester of (Enzyme free) b STR ibuprofen) Entrapment by TSP-EMR (organic) 160 cross-flow (aqueous) 400 filtration Lipase (Methyl ester (Enzyme free) STR of naproxen) Entrapment by TSP-EMR (organic) 300 cross-flow (aqueous) 300 filtration Fumarase STR (fumaric acid) (Enzyme free) ¢
4.50
1.12
Low
1.2x10 ~
High
2
Low 5.3x10 "3 8x10 "4
High
290
Low
Entrapment by UF-EMR 140 2375 High cross-flow filtration d aKm = 9 m m o l ' l "1, V m a x = 20 k + 2 - 6.86 s'l; aKm = 47 m m o l ' l l , V m a x = 4 mmol.1-1 h -1, k+2 = 1.37 s l ; aKin = 212 m M retool.l-l; V m a x = 9000 retool.1 -l h l ; k+2 = 5x1012 s "l, dk+2 = 8x1012 s -1
In addition to the membrane material, it is very important also the amount of enzyme immobilized [26, 55, 65, 66]. It is generally observed that the performance of immobilized enzyme as a function of density of enzyme on membrane (gprotein/Cm3) has an optimum. Increasing the gprotein/Cm3 the catalytic activity can be stable or increase until it reaches a maximum, after which, for further increase of protein loading no increase of activity is achieved. The initial increase of activity with the enzyme loading, indicates that the interaction with the support destabilizes the protein, and as the protein amount is increased it is more protected, and then more active. The constant activity is mainly due to mass transport limitations. It some cases, further increase of protein can even lead to decrease of activity. This happens when the protein-protein interactions
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at high concentration cause denaturation of protein conformation; either because of bound with the active site or because of proteolitic activity present in the crude enzyme preparation. The behavior of some enzyme as a function of loading amount is summarized in Figure 1, from which it can be seen that the same type of immobilized enzyme can show different behavior with different substrates. For
Activity
h.._
Enzyme loading
v
Figure 1. Qualitative behaviour of activity of immobilized enzymes as a function of enzyme loading: •
kipase C. rugosa in polyamide membrane (hydrolysis activity with olive oil) [26]
II
Lipase C. rugosa in polyamide membrane (hydrolysis activity with naproxen ester) [35] Lipase C. antarctica on Accurel0a)EP100 (Esterification activity) [55]
O
Lipase Humicola sp on Accurelc.)EP100 (Esterification activity) [55]
r]
Fumarase from porcine heart (hydrolysis activity)[ 19]
example, lipase from C. rugosa when used with triglycerides as substrate, shows constant behavior for initial increase of immobilized enzyme and after a decrease; while, when used with naproxen ester for initial increase of immobilized enzyme the activity increases and after, for further increase of enzyme, the activity decreases. This implies that the same enzyme conformation, achieved with the same type of immobilization, has different affinity for the different substrates. The use of proper conditions may allow to reduce the amount of enzyme with the result of reducing the mass transport limitations and cost.
2.1.2. The influence of hydrodynamics conditions In order to measure the kinetic properties of immobilized systems, it is necessary to work in conditions of reaction limited regime, where the mass transport rate is much higher compared to the reaction rate [67].
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Membrane Bioreactors For Biotechnology And Medical Applications - Drioli
Among the various methods to control the mass transfer coefficient, the fluid dynamics conditions, such as axial velocity and transmembrane pressure, may help in achieving the proper mass transfer properties. In two-separate phase (or biphasic) membrane reactors (TSP-EMR), where the mass transport occurs mainly by diffusion, the effect of the flow rate of organic phase (which contains the substrate) can be very different from that of the aqueous phase (which extracts the product) on the reactor performance. An example of the mentioned behavior is reported in Figure 2 and refers to the lipase using olive oil and cyano ester of ibuprofen as substrate. For the studied values, the observed reaction rate decreases with increase of the organic flow rate, for both substrates. Vice versa, the observed reaction rate increases with increasing the aqueous flow rate. This behavior is due to the fact that the higher transport of substrate (achieved at higher axial velocity of the organic phase) a higher reaction rate is obtained. On the other hand, the fixed axial velocity of the aqueous phase was not able to promote the extraction of the reaction product, which therefore, inhibited the reactor performance. This show the importance of both transport of substrate and product through the enzyme-loaded membrane 0.14 0.12 .,,.... 0.1
.,=
0.08 E g 0.06
l e Olive oil &CNE
0.04 0.02 0 0
1O0
200
300
Qa (mmol/I h)
Figure 2. Effect of axial flow rate of organic and aqueous phases on the performance of two-separate phase enzyme membrane reactor
reactor. In ultrafiltration enzyme membrane reactors (UF-EMR), where the mass transport occurs by convection, the permeate velocity is the parameter that mostly affects the reactor performance, as it influences the residence time of reagents within the enzyme-loaded membrane. The residence time, T, is calculated as the ratio between the membrane thickness, L, and the permeate velocity, L/T. Figure 3 shows the influence of residence time on the conversion of fumaric acid into L-malic acid using fumarase immobilized in a ultrafiltration membrane reactor [19]. The effect of the amount of enzyme loading is also illustrated.
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- Drioli
0.9 0.8 .-. 0.7
~
n-n
~o 0.6 •E
0.5 0.4 0.3 • 0.3 mg/cm 3
0.2 0.1
• 1.14 mg/cm3
0 0
i
i
i
~
i
100
200
300
400
500
600
Residence time (s) Figure 3. Influence of residence time on the efficiency of and ultrafiltration enzyme membrane reactors
2.1.3. The influence of reaction microenvironment on enzyme enantioselectivity Thanks to their properties, lipases have found applications in various fields of biotransformations. One of the most recent and attractive field is the synthesis of chiral compounds. The importance of using optical pure enantiomers became clear in the recent years thanks to the understanding of the different effects of enantiomers, where often one has the positive properties whilst the other can cause serious side effects [68]. The chiral market represents about the 25% of the drug market; on 1993 only the 3% of this was represented by pure enantiomers while the rest was sealed as racemic mixtures; on 1997 more than 75% of chiral drugs were marketed as pure enantiomers. This picture shows the rapid change towards the use of pure enantiomers in the pharmaceutical field, that had his first push in 1992 when FDA issued a policy statement. The same trend is expected to happen for food and agro-chemicals. On the basis of this context, many efforts are focused on the development of technologies to produce optically pure isomers. Although the most common method to produce enantiomers at industrial level is the chemical synthesis of a racemic mixtures followed by separation of the two enantiomers by diastereomeric crystallization, new resolution processes are being considered. Among others, membrane processes are today investigated for this purpose. On the basis of the current literature, there are two schemes on the use of membrane technology to produce enantiomers [69]. In one case, the membrane itself is intrinsically enantioselective and separates the wonted isomer on the basis of the pore spatial conformation. In the other case, a kinetic resolution catalyzed by an enantiospecific catalyst is combined with a membrane 193
Membrane Bioreactors For Biotechnology And Medical Applications - Drioli
separation process: the chiral system is represented by the catalyst, the membrane only separates the product from the substrate on the basis of their physical chemical properties. Studies on the use of lipase-catalyzed enantioselective biotransformations in multiphasic membrane reactors, for the production of pure (S)-naproxen have been carried out at our laboratory. Particular attention was focused on the parameters that influence the catalytic properties of the lipase immobilized on polymeric membrane. The performance of lipase immobilized with the same method on different type of membranes, which were then used in different reactor systems, has been investigated. The lipase was immobilized by cross-flow ultrafiltration, so that to obtain a stable irreversible lipase fouling within the sponge layer or on the lumen surface of asymmetric polymeric membranes. The lipase-loaded membranes were used to carry out the kinetic resolution of (R,S)-naproxen methyl ester to obtain the pure (S)-naproxen acid, in twoseparate phase enzyme membrane reactor (TSP-EMR) and in emulsion enzyme membrane reactor (E-EMR) systems. In the TSP-EMR the organic and aqueous phases are separated by the enzyme-loaded membrane and the mass transport through the membrane occurs by diffusion. In this system, it is difficult to obtain the interface at the level where the enzyme is immobilized. Therefore, the observed catalytic activity and enantioselectivity resulted lower compared to the free native lipase, because they were negatively affected by the fact that some enzyme worked in organic phase, some in aqueous phase and only part worked at the organic/water interface. Polyamide membranes allowed to obtain higher enantioselectivity compared to polysulphone membrane, and in general, for both types of membranes, the sponge layer resulted a better immobilization site, compared to the thin selective layer, due to its higher loading capacity. The catalytic stability of immobilized lipase was very good for both type of membranes [70]. In the E-EMR an organic in water (O/W) microemulsion, prepared by membrane technology, with a size distribution from 1 to 4 ~m, was fed to the lipase-loaded membrane from shell to lumen, so that pores into the sponge layer were filled with the microemulsion. The thin selective layer allowed only water to permeate through the membrane whilst the organic phase was retained. Therefore, while the substrate is kept in contact with the immobilized lipase, the product is removed in the permeated water. In this system mass transfer through the membrane occurs by convection, and between the two phases it occurs by diffusion. The great advantage of the E-EMR is that all the immobilized enzyme is able to work at the organic/water interface thanks to the presence of the microemulsion within the pores. This allows to achieve catalytic activity and enantioselectivity for the immobilized lipase comparable to the free native lipase, and much higher stability, with overcome of the product inhibition
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Membrane Bioreactors For Biotechnology And Medical Applications - Drioli
effects. In this system, each finger of the sponge layer works as a continuous tank emulsion microreactor. The experiments with E-EMR demonstrated that the lower enantioselectivity of immobilized lipase observed in TSP-EMR is due to microenvironment reaction conditions rather than to deactivation of the enzyme due to immobilization. In fact, for comparable amount of enzyme loading, immobilized with the same procedure on the same type of membrane support, the E-EMR showed about 30% higher enantioexcess compared to the TSPEMR.
2.2
Applications Table 2 summarizes some of the applications of immobilized enzymes at large scale. Table 2. Examples of industrial application Purpose Ref. Biocatalyst Immobilization F,scherichia coli Entrapped in Production of L-aspartic acid 56 cells polyacrylamide Entrapped into the Hydrolysis of the milk and whey lactose 57 Lactase (13galactosidase) and fibres of cellulose acetate used for the Synthesis of the dipeptide AspartameTM 58 Termolysin Pseudomonas Immobilized with Production of L-alanine 59 dacunahe glutaraldehyde Glucose-isomerase Reticulate with Production of fructose concentrated 60 glutaraldehide syrups Amino acylase Immobilized on Production of L-amino acids from 61 DEAE-Sephadex racemic mixtures Brevibacterium Entrapped in Production of L-malic acid 62 ammoniagenes polyacrylamide Lipase Entrapped in hollow- Production of (2R,3S)-trans isomer of 49 fiber asymmetric methyl ester of 4-methoxyphenylglycidic membranes acid The major technological difficulties in using biological immobilized systems on an industrial level are related with the life-time of the enzyme; the availability of pure enzyme at an acceptable cost; the necessity for biocatalysts to operate at low substrate concentrations; microbial contamination; the lack of general assumption and the consequent need for individual development on a trial and error basis. Nevertheless, nowadays some new conditions (such as the higher attention of governments towards environment and health care; the need of innovative technologies for a sustainable growth, etc.) may force the industry to develop new, safe, clean, and low energy consumption technologies such as the membrane bioreactors. .....
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Membrane Bioreactors For Biotechnology And Medical Applications - Drioli
Membrane technologies used in the production of optical pure isomers have been recently described [69, 71]. The major problems in the production of these compounds on a large scale concern: the requirements for expensive cofactors; the low water-solubility of the substrates; the separation and purification of the products from complex solutions. Studies carried out mainly on a laboratory scale indicate that in many cases the use of the appropriate type of membrane and membrane reactor design can overcome these difficulties. For example, UF-charged membrane reactors can be applied to retain cofactors [72]; two separate phase membrane reactors can be used for the bioconversion of low water-soluble substrates [49]. In addition to the use of enzymes in organicaqueous systems, their use in pure organic solvent is increasing. Enzymes in organic media are able to work in microenvironments that contain very little quantifies of water (usually less than the solubility limit). Immobilized enzymes operating in organic media show novel properties, such as enhanced stability and altered substrate specificity. An investigation on transesterification reaction in anhydrous tetrahydrofuran carried out with free and immobilized lipase demonstrated that the enzyme immobilized on zirconia membranes is more active compared with the enzyme in suspension and that the selectivity towards a reaction intermediate changed, resulting in the production of a non-naturally available flavanoid [54]. In many instances, when the product is obtained by fermentation, it is present as a component of a complex solution from which it needs to be separated and purified. In these cases, integrated membrane systems can be used for continuous production and downstream separation. For example, the production of L-lactic acid is obtained by continuous fermentation in a membrane fermentor. This consists of a traditional fermentor combined with an ultrafiltration unit. The solution recovered as permeate contains the product, Llactic acid, together with other small molecules that are not retained by the membrane, whilst the cells and macromolecules are recycled back to the bioreactor. The purification of the product can be achieved by membrane-based solvent extraction carried out through two membrane contactors. Electrodyalisis [73], nanofiltration and reverse osmosis [74, 75], membrane crystallization [76] are other separation processes that can be used in the downstream processing. In recent years, membrane technology has also attracted considerable attention for the treatment of waste-water [77,78]. The reason of this lies on the increase of potable water demand and on the increase of domestic and industrial wastewater discharges linked to the population growth. In addition, more restricted legislation towards environmental care, forces industries to minimize the input of energy and water. Two different membrane bioreactor configurations have been mainly investigated: MBRs with the membrane module outside the activated sludge tank, and MBRs with membrane submerged in the biological reactors. Due to
196
Membrane Bioreactors For Biotechnology And Medical Applications - Drioli
lower energy requirements, the submerged membrane bioreactors have found more success than the external ones. Membranes and membrane bioreactors have been studied for the removal of nitrogen from wastewater [79, 80]. Zoh et al. at University of California, combined a membrane bioreactor to a ceramic cross-flow ultrafiltration module for treating a synthetic wastewater containing hydrolysis byproducts of high explosive RDX (Hexahydro-l,3,5trinitro- 1,3,5-triazine) compound [81 ]. The state-of-the-art of submerged membranes has been recently depicted by R. Ben Aim who has discussed the potentialities (low energy input, long term operation without cleaning, less dependence on variation of rheology behavior with concentration) and limits (high membrane area due to low fluxes linked to the small transmembrane pressure) of the technology. Particular attention was devoted to the use of submerged membranes associated with aeration and on the effect of air bubble size on fouling [82]. Many studies in the field are oriented to the investigation of operating conditions and optimization, such as the effect of air sparging on cell viability; the nitrification rate and COD removal; the use of low cost membranes such as non-woven polypropylene to reduce capital costs; the conversion of organic wastes to energy and fertilizers. Others aim at showing the application at industrial level: Tazi-Pain reported the results of the first year operation of the BIOSEP (R) process, a full scale sequenced aeration MBR for municipal wastewater treatment. He claimed that this is the first plant in the world capable to achieve, in a single tank, the biodegradation of the organic matter, the nitrification-denitrification reactions and the physicochemical elimination of the phosphates [83]. The plant uses an immersed membrane system equipped with MF/UF hollow fibers supplied by Zenon Environmental Inc. Some examples of membrane bioreactors for waste water treatment at large scale are summarized in Table 3. 3.
MEMBRANE BIOREACTORS IN MEDICAL APPLICATIONS
Polymeric semipermeable membranes and membrane processes play a pivotal role in replacement therapy for acute and chronic organ failure and in the management of immunological disease. All clinical blood purification methods employ membrane devices [2].The next generation of artificial organs and tissue therapies is almost certain to be similarly grounded in membrane technology. Membranes of suitable molecular weight cut-off are used in bioartificial organs (e.g., pancreas, liver) using isolated cells, as selective barriers to prevent immune system components from getting into contact with the implant, while allowing nutrients and metabolites to permeate freely to and from cells. The use of membranes in bioartificial substitutes and tissue engineering date back to the 197
Membrane Bioreactors For Biotechnology And Medical Applications - Drioli
Table 3. Membrane bioreactors in water treatment at large scale application
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year 1933, when Vincenzo Bisceglie in Bail Italy encased mouse tumor cells in a nitro-cellulose membrane and inserted them into the abdominal cavity of guinea pig, to show that cells were not killed by an immune reaction in the pig [84]. Subsequently many researchers focused on the development of immunoprotective membranes to prolong the life of transplant. Currently two important areas of interest are the bioartificial pancreas for the treatment of insulin-dependent diabetes and the liver assist device for the temporary treatment of acute liver failure. Membrane capsules containing dopamine secreting cells also are being explored for treating of Parkinson's disease, a progressive brain disorder characterized by a deficiency of the neurotransmitter dopamine [85]. Immunoprotective membrane cell transplants are being investigated to treat other nervous system disorders. Polymer membranes also are being explored to block cell adhesion or scar tissue formation, for example after surgery, and thus improve wound healing. In addition, the membranes are being investigated for prevention of restenosis (coronary artery narrowing) after angioplasty [84]. In membrane bioartificial organs, cells are compartmentalized by means of semipermeable membranes that permit the transport of nutrients and metabolites to cells and the transport of catabolites and specific metabolic products to blood. The membrane must avoid the contact between xenogenic cells and patient's blood to prevent immunological response and rejection of xenograft. Membranes act as means for cell oxygenation and in the case of
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anchorage-dependent cells as substrata for cell attachment and culture. As a result, the type of membrane to use in a bioartificial organ must be chosen on the basis of its permeability characteristics as well as on its physico-chemical properties related to the separation process. 3.1.
Membrane bioartificial pancreas
In patients with insulin-dependent diabetes, the cells that normally produce insulin don't function, so patients must receive insulin injections to regulate their blood glucose level. This approach in the long term result in the so-called "diabetic complications" such as kidney failure, neuropathy, which lead to the patient's severe incapacity or death. Transplantation of the whole pancreas, pancreatic tissue fragments or islets of Langerhans would assure the required glucose-related insulin secretion, but the rejection of implants and scarce availability of donor organs limits this therapeutic approach [86]. An alternative approach is the development of membrane bioartificial pancreas using isolated islets of Langerhans or single beta-cells, which are able of sensing the plasmatic glucose concentration and produce insulin amounts related to the actual glycemia, entrapped by means of membranes. In 1970 W.L. Chick and colleagues transplanted isolated islets protected by hollow-fibre ultrafiltration membrane (an acrylonitrile-vinyl chloride copolymer) into dogs made diabetic by surgically removing the pancreas. The device consists of a chamber through which passes a copolymer membrane connected to standard vascular grafts. Islets are placed inside the chamber, through ports in the housing into the cavity, but are outside of the blood stream. Nominal molecular porosity of 80,000 Dalton permits free diffusion of nutrients and insulin across the membrane but inhibits the entry of immunoglobulins and immunocytes from the blood stream into the chamber (Figure 4a) [87]. Since 1970 research efforts were devoted to the development of a hybrid bioartificial membrane pancreas [84-87]. Different bioartificial pancreas were designed in four physical types: hollow fibers, capsule, coatings and sheet. The many different types of prosthetic devices proposed so far can be grouped in three main categories" extravascular devices, intravascular devices and microencapsulated islets of Langerhans [89-98]. In the first case, the tissue is enclosed between membranes, if in a flat sheet configuration, or in the lumen of hollow fibre membranes, and then implanted in an extravascular site (see Figure 4). Other researchers [88] proposed hollow fiber device as extravascular bioartificial pancreas. This device consists of hollow fiber with 0.5 mm diameter, 20 mm length, containing 80,000 islets/ml (Figure 4b). The extravascular systems generally suffer from an intrinsically slow insulin response following changes of blood glucose concentration, limited by the purely diffusive mass transport and by the fibroblastic response of the host.
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waste products from them. Normoglycemia has been reported after intraperitoneal transplantation of alginate-polylysine microencapsulated allogenic and xenogenic islets in diabetic animal models and recently also in men. However, graft survival is always limited to several weeks. Graft failure is interpreted as non-specific immune response i.e., foreign body reaction against the microcapsules resulting in a progressive overgrowth of the capsule and subsequent necrosis of islets. Researchers focused study on highly purified alginate and other biochemicals more biocompatible [91] and on novel membranes able to prevent permeation of low molecular weight humoral molecules released by xenogenic islets. The Islet Sheet is a thin planar bioartificial pancreas, licensed from Islet Sheet Medical Company, contains live, functional islets in an artificial polymer matrix (Figure 4h) [93]. Each sheet is several cm in diameter and contains 2 to 3 million cells and contains islets microencapsulated within in a mesh to increase the physical strength between two layer of semipermeable alginate membrane. A semipermeable membrane such as 0.2 ~tm cellulose ester filter membrane is saturated with crosslinking solution. The cells 4 to 6 sheets contain enough islet tissue to cure diabetes in an adult. The sheet is so thin (the overall thickness is 250 ~tm) that diffusion alone allows sufficient nutrients to reach the center of the sheet. A coat on the exterior of the sheet prevents contact between the cells inside and immune effector cells of the host as well as inhibiting diffusion of antibody and complement. The alginate membranes show high permeability to different solutes and excluded immunocompetent species. No immune suppression drugs are needed. The sheet may be removed or replaced at any time. Experiments on islet sheet began at the University of Chicago in September 1998, and moved to the University of Cincinnati and the University of Alberta in 2000. Large animal studies given encouraging results. The implantation of islet sheet in omentum of pancreatectomised dog permitted to return blood sugar to normal and at 60 days the blood sugar was lower at every measurement time [94]. 3.1.2. Membrane bioartificial liver
Each year in the US, approximately 150,000 people are hospitalized with liver disease and over 43,000 people die from it [95]. Transplantation, the only effective means of treating liver failure, is not an option for many patients. Ironically, the liver is a highly regenerative organ [96]. Supportive therapies as dialysis, hemofiltration, hemoperfusion for patients with acute liver failure have been proven ineffective because physical methods are not sufficient for the management of severe biochemical disorders. The development of an extracorporeal liver assist device, using isolated liver cells, to which a patient would be temporarily connected until he/she recovered or received a liver transplant could be a promising approach for 202
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patients with liver failure. Since fulminant liver failure is potentially reversible, the extracorporeal bridging of liver function would also be beneficial until the patient's own liver resumes functional activity.
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The first clinical report of bioartificial membrane liver was released in 1987 [97]. This device consisted of a hepatocyte suspension that was separated from the patient's blood by a cellulose acetate dialysis membrane. Since different devices have been proposed (Figure 5). These devices differ from one other for the type of material used to construct the bioreactor: membrane, glass or non-woven fabrics; for configuration: fiat or hollow fiber; for coating used: matrigel, collagen; for cell culture technique" microcarrier, spheroids, aggregates, etc, and for cell capacity. In devices using hollow-fibre membranes, isolated hepatocytes are usually loaded outside of the hollow fibres in the extracapillary compartment, while blood, plasma, or culture medium flow through the lumen of the hollow fibres (Figure 5b). In devices using fiat-sheet membranes, hepatocytes are loaded between fiat membranes in a sandwich way,
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while blood or culture medium flow outside of the membranes as shown in Figure 5c. Cells may be free in suspension, attached to walls or attached to microcarrier (Figure 5d). A key issue concerning the development of bioartificial liver is the maintenance of long-term viability and functions of hepatocytes: oxygen transport resistance and catabolite accumulation may limit the hepatocyte viability and metabolism [98-99]. To overcome the problem of maintenance of viability and liver specific functions in vitro of liver cells, Sussmann [100] used immortalized liver cell line in the extracorporeal hollow fibre device as biological element. This device, developed as Hepatix/Baylor unit, consisted of 10,000 individual hollow fibres with cut-off of 70 kDa. This has been tested on 11 patients with fulminant hepatic failure: four patients were sustained until transplant and two survived without the need of transplant and five died for complications. A disadvantage of this unit is the potential of seeding tumor cells into patients. Thus, other researcher proposed new culture models to address the problem of long-term maintenance of liver specific functions of liver cells inside of the bioreactor. Flendrig et al. [101], developed a bioreactor based on a spirally wound non-woven polyester matrix, three-dimensional framework for hepatocyte immobilization and hydrophobic polypropylene membranes for decentralized oxygen supply and CO2 removal (Figure 5e). Medium or plasma is perfused through the extrafibre space and in direct hepatocyte contact. Generally many investigators inoculate cells in the extrafibre space. In Nyberg's design hepatocytes are cultured in the hollow fibre lumens [102]. In this device hepatocytes are entrapped in a three-dimensional gel matrix and the extrafibre space of the bioreactor is perfused for 24 h with recirculating medium, as is shown in Figure 5f. The blood is passing through the extrafibre compartment and the gel-entrapped cells are perfused by means of medium to improve the supply of oxygen, nutrients, and the removal of catabolites. This device is scaled-up by using hepatocyte spheroids inside polysulfone membranes at University of Minnesota and recently was licensed to Algenix Inc. and it was approved by FDA for Phase I human clinical trials. An interesting device was designed by Gerlach et al. [103] by mixing different membrane fibres in various directions, thus improving the oxygen and nutrients supply to hepatocytes located in between such a network. The Figure 5g shows the scheme of such device licensed to Hybrid Organ GmbH. Four separate capillary membranes with different functions are utilized: plasma inflow by polyamide membrane; plasma outflow by polysulfone membranes; decentralized oxygen supply and carbon dioxide removal with low gradients by polypropylene membranes; sinusoidal endothelial coculture by hydrophilic membrane. This device was used for preliminary clinical studies. Based on dialysis model, the liver support system developed by Demetriou et al. [ 104] of Cedar Sinai Medical Centre of Los Angeles, consists 205
Membrane
Bioreactors For Biotechnology And Medical A p p l i c a t i o n s - Drioli
of a bioartificial liver cartridge that contains billions of pig liver cells and a machine that controls the flow of blood through the cartridge (Figure 5d). The commercial name of the device is "HepatAssist System" of Circe Biomedical and includes cellulose nitrate/cellulose acetate hollow fibre bioreactor containing porcine hepatocytes attached to a collagen-coated microcarriers, two charcoal columns, a membrane oxygenator, and a pump. Plasma from a patient flows through the cartridge. Pig liver cells on the outer surface of the fibres are in contact with the plasma through large pores on the fibre surface. This device was used in USA for a treatment of 6-7 h of 25 patients with severe acute liver failure. These reports encourage the further studies to address the problems related to the development of a bioartificial liver. In fact, although several groups worldwide are creating systems there is not still an ideal device. This is not surprising, given the complexity of liver functions. Liver support system must supply various liver specific functions including synthetic and detoxification activities for a time sufficient to allow recovery of a patient or maintenance of patients until transplantation. Recently, we evaluated in vitro the performance of a full-scale fiat membrane bioreactor (FMB) developed by Bader et al., that permits the culture of liver cells under in vivo-like conditions and at high-density culture [105-109]. In such bioreactor porcine hepatocytes are cultured within extracellular matrix between oxygen permeable fiat-sheet membranes. Isolated liver cells are located at a distance of 10-20 ~tm of extracellular matrix. This bioreactor provides culture conditions that improve liver specific functions of liver cells in vitro. The FMB is able to provide an in vivo-like microenvironment for liver cells: hepatocytes are arranged as a plate in 3-D coculture with intermingled nonparenchymal cells. In contrast to other bioreactors the FMB is based on the organization of liver cells as a plate within extracellular matrix in which each individual hepatocyte has its own membrane support and thereby its own oxygen supply position. In vitro studies demonstrated that the performance of a scale-up FMB using porcine hepatocytes is stable over a period of about 3 weeks and compares well with that of other systems present in literature. Isolated hepatocytes cultured in the FMB reconstitute many of the features of the liver in vivo. The cell concentration inside of the FMB increased in the first days of culture and then remained constant until 18 days. Specific metabolic functions of hepatocytes in terms of albumin synthesis, ammonia elimination and urea synthesis are sustained for the investigated culture time demonstrating thus the long-term maintenance of functional integrity of hepatocytes cultured in the FMB.
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3.1.3.~Cytocompatibility of membranes in bioartificial organs In membrane bioartificial organs using isolated cells as biological component, semipermeable membranes play more functions: they act as immunoselective barriers, as means for cell oxygenation and provide a large area for cell attachment. All these functions are important for the maintenance of cell viability and specific functions. In our experimental study we demonstrated that isolated rat liver cells cultured on oxygen-permeable membranes reconstitute many feature of the in vivo [98]. Cells cultured on polythetrafluoroethylene oxygen-permeable membranes maintained a morphological appearance of hepatocytes similar to their in vivo appearance and exhibited at high levels tissue specific functions in vitro in terms of albumin secretion, urea synthesis and drug biotransformation functions. This finding is of relevance for oxygen supply to cells in vitro both in batch and in large-scale bioreactor systems. In a membrane bioartificial organ, cells come into contact with the membrane surface. Therefore, the response of the biological components depends on surface properties of the used membrane. Physico-chemical properties including surface composition, surface charge, surface energy, and surface morphology, may affect cell adhesion and behavior. Recently, studies performed on isolated hepatocytes cultured on semipermeable polymeric membranes have indicated that wettable and rougher surfaces enhance adhesion and metabolism of isolated hepatocytes [ 110-112]. In particular, measurements of the wettability of membrane, expressed by the contact angle in the presence of different liquids permitted to evaluate and to compare surface free energy components of membranes with different physicochemical properties. Such measurements before and after modification of native membranes in culture medium resulted to be predictive index of their cytocompatibility and/or tissue biocompatibility. Therefore a material surface treatment might enable the adaptation of its surface free energy to biological requirements. The adhesion and activity of liver cells are affected by hydrophilic properties of membranes: surface free energy and its components. Isolated hepatocytes formed on membranes with high base parameters of surface free energy, such as cellulose acetate, polycarbonayte membrane, large cellular aggregates while on hydrophobic membranes such as polypropylene membranes cells spread to a large extent and the edge of cells were often indistinguishable in a network formed by fibrous proteins of extracellular matrix. For each membrane we observed that cell adhesion increased with increasing base parameter of membrane surface tension. In Figure 6 are reported the cell adhesion after 48 h of culture on membranes with different base parameter of surface tension. Liver cells adhere strongly on membranes with high value of base parameter. The absolute value of cell adhesion is higher in the presence of
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serum proteins adsorbed on the membrane surface, which change the wettability by increasing base parameter of surface tension. Adsorption of proteins from surrounding liquids is one of the key steps for the cytocompatibility of the membrane. We observed that the extent of serum protein adsorption is relatively large on hydrophobic membranes and relatively small on more hydrophilic membranes [ 112].
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The base parameter of surface free energy affected the metabolic functions of liver cells. We found a correlation between urea synthesis and base parameter of membrane surface tension. For each investigated membrane cells synthesise urea with a rate that increases with increasing of base parameter value of membrane surface tension (Figure 7). The metabolic activity is particularly expressed at high levels when cells are cultured on PC and CA membranes. Figure 7 shows that on CA and PC membranes the ability of liver cells to synthesize urea increases linearly with increasing the value of base parameter of surface tension and reached the value of about 120 mg/h in correspondence of the surfaces with respectively 46 mJ/m 2 for CA membrane and 54 mJ/m 2 for PC membrane. On such membrane surfaces cells exhibited rates of urea synthesis that significantly increase with time. These results suggest that there is a marked
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independently on the type of native polymeric membranes it is possible to improve cell adhesion and specific functions by changing their surface tension and their components. On the basis of our results, CA and PC membranes resulted to be good substrata for liver cell culture in vitro promoting cell adhesion and specific functions in term of urea synthesis. To this purpose, researches on interactions of cells with novel membranes made in our laboratory are running in order to develop more cytocompatible and biocompatible membranes.
3.1.4. Membrane biocompatibility When blood comes in contact with foreign substances by extracorporeal circulation, reaction to foreign body occurs in the blood. Thus, to evaluate the safety of the material is used a reaction to foreign body as an indicator of "biocompatibility". A biocompatible membrane is a membrane more affinous is to living body and that gives less reaction to foreign body.
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Materials in blood contact often initiate coagulation processes, affect platelet morphology and function and lead to an immunologic answer of the body. The first event of a complicated sequence that leads to thrombus formation and immunoreactions is the protein adsorption, which plays a mediating role in bioaccumulation, systemic foreign body reactions, tissue regeneration and is, therefore, crucial for biocompatibility and performance of medical devices. Thus, the development of biomedical polymers has to comprise the detailed investigation of protein adsorption. Several studies indicated that platelet activation increases in the positively charged membrane and decreases in the membrane having a microdomain structure in which hydrophilic groups and hydrophobic groups coexist randomly as molecules. Also the complement activation, which in the case of biomaterials proceeds via alternative pathway, is manly noted with cellulose membranes; free hydroxyl groups on the membrane surface is bonded with C3b and further with factor B which promotes the activation. Improvements of biocompatibility of cellulose membrane can be obtained by surface modification devoted to graft for example alkyl group to the hydroxyl group. Considering that functions of biointerfaces are mainly based on principles of molecular recognition, several powerful methods of surface functionalization including plasma treatments, grafting of reactive chains, coating of functional layers and others can be used to engineering news modified surfaces [113-114]. Furthermore, in addition to original cellulose membrane, a variety of synthetic polymers less potent in complement activation, can be used clinically. 30 ,.--, tN
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In this respect, study on biocompatibility of commercial and novel polymeric membranes made in our laboratory are running. Preliminary results show that human plasma fibrinogen and immunoglobulin G adsorb to a greatest extent on cellulose acetate membranes. Differently commercial membranes such as polypropylene and novel membranes such as modified polyurethane and polyetherketone membranes adsorb fibrinogen and immunoglobulin G to a leatest extent (Figure 8). On the basis of these preliminary results these membranes seem to have antithrombogenic properties. Further investigations are needed. 4.
CONCLUSIONS
The progresses made in the last years in membrane engineering and in the understanding of basic mechanisms for membrane transport phenomena, have a significant fall out not only in the industrial separation processes, in desalination, in gas treatments but also in biotechnology and in biomedical engineering. The potential advantages of membrane technology over more conventional approaches, include higher efficiency and reduced costs owing to the integration of bioconversion and product purification, thus reducing equipment costs and processing steps as today requested by a process intensification strategy. Enzymatic membranes is also contributing to the growth of new research area, such as nonaqueous enzymology, the use of antibodies as highly specificcatalysts, use of biomimetic catalysts (such as cyclodextrines) and the development of novel biosensor for diagnostic purposes. The results reached in the development of biocatalytic membrane reactors for the industrial applications and for the realization of artificial organs are important examples. Further progresses will be realized by the integration of these systems, by the study of new more selective and resistant organic and inorganic membranes, by their further miniaturization. REFERENCES [ 1]
[2]
[4]
Malmsten M. Protein adsorption at the solid-liquid interface, In Protein Architecture: interfacing, molecular assemblies and immobilization biotechnology Yuri Lvov and Helmuth Mohwald (eds.), pp. 1-24, Marcel Dekker Publisher, USA, 2000 De Bartolo L and Drioli E. Membranes in artificial organs. In Biomedical and Health Research vol. 16: New Biomedical Materials - Basic and Applied Studies, Haris PI and Chapman D, (eds), pp. 167-181, IOS Press: Amsterdam/Berlin/TokjoAVashington, 1998 Takamatsu S et al. Production of L-alanine from ammonium fumarate using immobilized microorganisms - elimination of side reaction Eur. J. Appl. Microbiol. Biotechnol. 1982; 15:147
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[5] [6]
[7]
[8] [9] [10]
[11]
[12]
[13] [14]
[15] [16]
[17]
[18] [19]
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New Insights into Membrane Science and Technology: Polymeric and Biofunctional Membranes D. Bhattacharyya and D.A. Butterfield (Editors) © 2003 Elsevier Science B.V. All rights reserved.
Chapter 10
Structural and performance characteristics of hemodialysis membranes D. Gao I and W. R. Clark2, 3 .
1Department of Mechanical Engineering/Center for Biomedical Engineering, University of Kentucky, Lexington, KY, and Department of Thermal Science & Energy Engineering, University of Science and Technology of China, Hefei, Anhui, China 2
Renal Division, Baxter Healthcare Corp., McGaw Park, IL
3Nephrology Division, Indiana University School of Medicine, Indianapolis, IN *Corresponding author, Phone: (317) 613-2315; Ext. 327, Fax: (317) 613-2317 e-mail: [email protected] ABSTRACT Recent trends show a progressive increase in the use of modified cellulosic and synthetic dialyzers and a corresponding decrease in the utilization rate of unmodified cellulosic dialyzers. The purpose of this paper is to describe current membrane and dialyzer technology, with the focus almost entirely on modified cellulosic and synthetic membranes. A general overview of membrane-related determinants of dialyzer performance is first presented, followed by a discussion of specific characteristics of some of the more commonly used membranes and dialyzers. 1.
INTRODUCTION
End stage renal disease (ESRD) has been a major medical problem worldwide. According to the United States Renal Data System (USRDS), in the United States alone for year 2000, there were 372,407 people projected for the ESRD, and the life of 281,355 of these ESRD patients was sustained on hemodialysis using an artificial kidney (or so-called hemodialyzer). The total cost of Medicare for ESRD patients was projected to be $14.2 billion in year 2000 in USA.
219
Structural And Performance Characteristics Of H e m o d i a l y s i s M e m b r a n e s - Clark
Dialysate out
Dialysate in
t !
Blood out
Blood in
--? ...__~
I . . . . . . . . . . . .
iIiij11s~slsl
/ /
Header or Manifold 11
.-"
Cross-Sectional View -/ i i
I
I
I
I
I
'otting Region Hollow Fibers
Figure 1. A general schematic diagram of an artificial kidney. An artificial kidney is a countercurrent mass exchanger containing approximately 10,000 hollow fibers made up of porous membranes. This mass exchanger has two compartments: the blood compartment (total internal space of all hollow fibers) and the dialysate compartment (space among the hollow fibers). During the hemodialysis, blood is running through the blood compartment, while the dialysate is flowing counter-currently through the dialysate compartment. The uremic toxins are removed from the blood into the dialysate through both diffusion and convection across the hollow fiber membranes.
In engineering perspectives, an artificial kidney is actually a countercurrent mass exchanger containing approximately 10,000 hollow fibers made up of porous membranes. The schematic diagram of an artificial kidney/dialyzer is shown in Figure 1. This mass exchanger has two compartments: the blood compartment (internal space of the hollow fibers) and the dialysate compartment (external space of the hollow fibers). As their names imply, the blood is introduced through the blood compartment, while the dialysate (cleansing salt-balanced water solution) is flowing counter-currently through the dialysate compartment. The porous membranes (hollow fibers) act as an interface where solute mass transfer between the blood and the dialysate 220
Structural And Performance Characteristics Of Hemodialysis Membranes - Clark
takes place. The overall mass transfer area of an artificial kidney is maximized by the use of these hollow fibers. The countercurrent blood and dialysate flows are introduced to further optimize mass transfer efficiency of the artificial kidney to remove uremic toxins in the blood and excess plasma water from patients suffered from ESRD. As discussed recently (1), the first hollow fiber artificial kidney used clinically in the 1960's consisted of a 1.0 m 2 unmodified cellulosic membrane. Hollow fiber dialyzers with highly permeable polyacrylonitrile (2) and polysulfone membranes (3) were subsequently developed, as were dialyzers with modified cellulosic membranes. Over time there has been a progressive increase in the use of modified cellulosic and synthetic dialyzers, predominantly in the high-efficiency and high-flux segments, respectively, and a corresponding decrease in the utilization rate of unmodified cellulosic dialyzers (4). In fact, these latter dialyzers play a relatively limited role in the current global hemodialysis (HD) market, for which approximately 90 million units are now manufactured annually. The movement away from unmodified cellulosic dialyzers has been driven largely by the desire to extend the molecular weight spectrum of solute removal and to minimize complement activation. In addition, retrospective data suggest outcomes of patients treated with these types of dialyzers are inferior to those achieved with modified cellulosic and synthetic dialyzers (5, 6). The purpose of this paper is to describe current membrane and dialyzer technology. A general overview of membrane-related determinants of dialyzer performance is first presented, followed by a discussion of specific characteristics of some of the more commonly used dialyzers. 2.
H O L L O W FIBER MEMBRANE CHARACTERISTICS
In the assessment of the clinical effects of a particular dialyzer, the membrane justifiably receives the most scrutiny. In this section, several important characteristics determining membrane performance are discussed. These include hollow fiber dimensions, surface area, porosity, and water permeability. 2.1.
Hollow Fiber Dimensions An individual hollow fiber can be viewed as a solid cylinder in which a central region has been removed ("cored out") to form the blood compartment. The process of manufacturing ("spinning") a hollow fiber membrane is a complex one incorporating aspects of a broad array of scientific disciplines, including polymer chemistry, thermodynamics, reaction kinetics, and chemical engineering (7). Although clear differences exist among the various types of
221
Structural A n d Performance Characteristics O f Hemodialysis Membranes - Clark
membranes (see below), a number of common features are evident. From a structural perspective, most hollow fibers have a relatively standard inner (blood compartment) diameter (approximately 180-220 ~tm) and length (approximately 20-24 cm). These parameters are dictated essentially by the operating conditions used during HD and are the result of a compromise between opposing forces. On one hand, a relatively small hollow fiber inner diameter is desirable because it provides a short diffusive distance for solute mass transfer. At a given blood flow rate, a lower inner diameter also provides a higher shear rate, resulting in greater attenuation of blood-side boundary layer effects (8). However, a decrease in hollow fiber inner diameter also has undesirable effects. Fluid flow along the length of a cylinder (i.e., the axial flow) in many situations is governed by the Hagen-Poisseuile equation (9): QB = AP / (8~tL / r~r4)
(1)
In this equation, QB is blood flow rate, AP is axial pressure drop, ~t is blood viscosity, L is fiber length, and r is hollow fiber radius. A specific application of this equation is axial blood flow (i.e., from the arterial to the venous end) in a hollow fiber membrane during HD. A more general form of Eq. 1 is: QB = AP / R
(2)
From Equations 1 and 2, the resistance to blood flow (R) is: R = 8gL / nr 4
(3)
Due to the inverse relationship between R and r4, a small decrease in hollow fiber inner diameter induces a large increase in flow resistance. Eq. 3 also demonstrates that increases in fiber length and hematocrit (~t) are associated with an increase in flow resistance. In turn, as indicated by Eq. 2, an increase in flow resistance results in an increase in axial pressure drop at a constant blood flow rate. In contemporary HD practice, patients with progressively rising hematocrits are treated commonly with large surface area dialyzers of high water permeability. These dialyzers require relatively high blood flow rates (at least 350 mL/min) to derive maximum benefit from a solute removal perspective. The high flow resistance and associated large axial pressure drop in this specific scenario result in significant backfiltration of dialysate into the blood compartment under normal HD operating conditions (10). The combination of significant backfiltration and contaminated dialysate increases the likelihood of cytokine-inducing substance transfer, as has been discussed recently (11). 222
Structural And Performance CharacteristicsOf HemodialysisMembranes -Clark
Modifications in considerations.
hollow
fiber
dimensions
are
constrained
by
these
2.2.
Surface Area For a particular hollow fiber, the inner annular surface represents the nominal blood compartment surface area and is the theoretically maximal area available for blood contact. For the entire group of fibers comprising a dialyzer, total nominal surface area then depends on fiber length, inner diameter, and overall number, the latter of which varies generally from approximately 7,000 to 14,000. A frequently asked question relates to the manner in which membrane surface area is calculated. For the inner annular region of a single hollow fiber described above, the surface area (A) is given by the equation defining the surface area of a cylinder: Afiber
=
2nrL
(4)
Based on assumed values of r = 100 gm (10 -4 m) and L = 24 cm (0.24 m), the surface area of an individual hollow fiber can be calculated as: Afiber
(2n)(10 -4 m)(0.24 m) = 1.51x10 -4 m 2 "-
For the large surface area dialyzers routinely used now, the total number of fibers (N) typically used is approximately 12,000. Therefore:
(5)
Adialyzer-- ( A f i b e r ) ( N )
= (1.51x10 -4 m2)(12,000) = 1.81 m 2 2.3.
Membrane Pore-Related Characteristics To provide rough quantitative estimates of pore-related parameters for a hollow fiber membrane used in HD, the straight cylindrical pore model can be used (12). As shown in Figure 2, this model assumes a membrane's pores all have the same radius (rp), which is larger than the radius (rs) of the hypothetical solute shown. In addition, the directional orientation of the cylinders is assumed to be perpendicular to the flow of blood and dialysate. As noted above, the Hagen-Poisseuile equation governs fluid flow through cylinders in many situations, applying also to transmembrane ultrafiltrate flow through the pores in this membrane model. As such, the rate of ultrafiltrate flow is directly related to the fourth-power of the pore radius (i.e., 1"4) at a constant transmembrane
223
Structural And Performance Characteristics Of Hemodialysis Membranes -Clark
pressure. Thus, although the number of pores also influences water permeability (i.e. hydraulic conductivity), the membrane characteristic that most directly influences water permeability is mean pore size. On the other hand, the diffusive properties of a dialysis membrane are determined mainly by the the porosity (pore density) and, to some extent, the pore size (13). Based on the cylindrical pore model described above, membrane porosity is directly proportional to both the number of pores and the square of the pore radius (r2). Therefore, diffusive permeability is also strongly dependent on pore size but not as strongly as is water permeability. The major pore-related determinants of flux (r4) and diffusive permeability (number of pores, r 2) differ sufficiently that the two properties can be independent of one another for a
1 --
oo!101° .
10~ 10= Pore Radius Rp (A) 2Rs--~
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O
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!
Figure 2: Pictorial representation of idealized, equal-sized pore membrane model (lower panel) and the associated relationship between pore size and solute permeability (upper panel). Reprinted with permission from (12) particular hemodialysis membrane. Specifically, cellulosic high-efficiency dialyzers typically have high diffusive permeability values for small solutes but low water permeability. On the other hand, there are examples of high-flux 224
Structural And Performance Characteristics O f Hemodialysis Membranes - Clark
dialyzers that have significantly lower small solute diffusive permeabilities than comparably sized dialyzers of much lower water permeability. It should be noted that HD membranes used in actual clinical practice demonstrate pores having a distribution of radii and tortuous (non-cylindrical) structures. As recently discussed (13), this pore size distribution and its influence on a membrane's sieving properties may differ significantly among the different dialysis membranes (Figure 3). The membrane represented by curve A has a large number of relatively small pores while the membrane represented by curve B has a large number of relatively large pores. Based on the relatively narrow pore size distributions, the solute sieving coefficient vs molecular weight profile for both membranes has the desirable sharp cut-off, similar to that of the native kidney. However, the molecular weight cut-off for membrane A (approximately 10 kDa) is
, .A
I0'
1.0
B
/", .~ a
!,
0.8
'
I
0
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0
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0
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20
0,0
~
30
40
50
60
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Figure 3: Pore size distribution and solute sieving coefficient profiles for three hypothetical dialysis membranes. The left diagram shows the number of pores as a function of mean pore size while the fight diagram shows the relationship between solute sieving coefficient and molecular weight. Reprinted with permission from (13).
consistent with a high-efficiency membrane while that of membrane B (approximately 60 kDa) is consistent with a high-flux membrane. In addition, primarily due to the large number of pores, both membranes would be expected to demonstrate favorable diffusive transport properties. On the other hand, membrane C exhibits a pore size distribution that is unfavorable from both a
225
Structural And Performance Characteristics Of Hemodialysis Membranes - Clark
diffusive transport and sieving perspective. The relatively small number of pores accounts for the poor diffusive properties. In addition, the broad distribution of pores explains not only the "early" drop-off in sieving coefficient at relatively low MW but also the "tail" effect at high MW. This latter phenomenon is highly undesirable as it may lead to unacceptably high albumin losses across the membrane. 3. NON-MEMBRANE-RELATED DETERMINANTS OF DIALYZER PERFORMANCE In the assessment of the clinical effects of a particular dialyzer, the membrane justifiably receives the most scrutiny. Although the membrane itself is a key determinant of overall dialyzer function, the manner in which the membrane interacts with other components of the dialyzer also is very important. In Figure 4, the sequence of events resulting in the conversion of polymethlymethacrylate (PMMA) hollow fibers to a dialyzer is shown (14). After removal of glycerin (used in the hollow fiber preparation), fibers are overed with spacer yams, which
t 4, [ R b e r bunclling i
~
p~,tl,n~
÷
1
I. . . . ÷
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......
Figure 4: Sequence of steps in the conversion of a hollow fiber to a dialyzer for PMMA. Reprinted with permission from (14).
are filaments designed to create optimal spacing between fibers (15). Subsequently, the fibers that eventually serve as the collective membrane in the dialyzer are assembled ("bundled") and inserted in the dialyzer casing. The
226
Structural And Performance Characteristics Of Hemodialysis Membranes - Clark
fiber bundle then is "potted" (encapsulated) at both ends with silicone rubber and cut. Finally, the dialyzer is placed in a pouch and the entire unit is sterilized, usually by ethylene oxide, gamma ray, or steam. Several of these manufacturing steps may have a direct effect on dialyzer function and performance. Fiber bundle configuration and spacing have a major impact on mass transfer, as has been demonstrated recently (16, 17). One consideration is the spacing of fibers within the bundle. However, of equal or more importance is the degree to which the dialyzer jacket is "packed" with fibers. A dialyzer's packing density is the ratio of the area comprised of fibers to the total area, based on a transverse cut through the dialyzer. Empirically, packing densities less than approximately 50% imply insufficient membrane surface area for an appropriate set of flow rates. On the other, values greater than approximately 60% are associated with a high risk of dialysate flow maldistribution in which dialysate is "channeled" to the peripheral aspect of the fiber bundle, at the expense of flow to the inner bundle area. Finally, sterilization technique may influence the dialyzer performance through an effect on mean membrane pore size. 4. DIALYZER CLASSIFICATION BY MEMBRANE COMPOSITION (TABLE 1)
4.1. Unmodified Cellulosic Dialyzers The constituent component of cellulosic membranes is ccllobiose, a saccharide found in a number of naturally occurring substances (18). From the perspective of blood's interaction with a cellulosic membrane, the most important characteristic of ccllobiose is its high density of hydroxyl groups. Although the contact of blood with any artificial surface elicits activation of the alternative complement pathway, the abundance of hydroxyl groups makes this phenomenon particularly pronounced for unmodified cellulosic dialyzers. This cellulosic characteristic was deemed clinically undesirable when first reported and has contributed to the progressive decline in unmodified cellulosic use over the years. However, the relatively long duration of popularity of cellulosic membranes can be explained largely by their particular suitability for a diffusion-based procedure like hemodialysis. The underlying hydrogel structure of these membranes and their tensile strength allow the combination of low wall thickness (see below) and high porosity to be attained in the fiber spinning process (19). These characteristics allow the attainment of high rates of diffusive membrane transport and efficient removal of small, water-soluble uremic solutes, such as urea and creatinine. Another characteristic feature of these membranes is symmetry with respect to composition, implying an
227
Structural And Performance Characteristics O f Hemodialysis Membranes - Clark
Table 1: HemodialysisMembranes Unmodified Cellulosic Modified Cellulosic (Subst. Group)
Svnthetic
Cuprophan
Cellulose (Di) Acetate (acetate)
Polysulfone
CupmmmoniumRayon
Cellulose Triacetate (acetate)
Polyamide
SCE
Hemophan (tertiary amine)
Polyethersulfone
SMC (benzyl)
PAN
Vitamin E-bonded
PMMA
essentially uniform resistance to mass transfer over the entire wall thickness. On the other hand, these membranes are characterized by low mean pore size and pronounced hydrophilicity, such that neither transmembrane nor adsorptive removal of middle and larger sized uremic toxins is significant typically (20, 21).
4.2.
Modified Cellulosic Dialyzers Similar to regenerated cellulose membranes, modified cellulosic membranes are characterized by low wall thickness values, typically in the 6 15 gm range, and symmetric structures. However, dialyzers composed of these membranes, first used for HD in the 1980s, cause less pronounced complement activation and generally have larger mean pore size (22) than their unmodified cellulosic counterparts. This latter characteristic results in higher water permeability and middle molecule clearances, relative to the unmodified cellulosic class. The two most commonly used modified cellulosic dialyzers contain membranes in which the hydroxyl replacement mechanisms are quite different. For cellulose acetate membranes (rigorously cellulose diacetate) (23), approximately 75% of the hydroxyl groups on the cellulosic backbone are replaced with an acetate group. As opposed to a hydroxyl group, an acetate group does not bind avidly to a C3 molecule to initiate activation of the complement cascade. Consequently, complement activation is attenuated, as is the leukopenic response, in which the WBC (white blood cell) count decrease from baseline is usually in the 35-40% range. Because production of cellulose triacetate membranes involves complete hydroxyl substitution replacement, further attenuation of complement activation and leukopenia is achieved (22). Along with cellulose acetate dialyzers, Hemophan® dialyzers are the most commonly used products containing modified cellulosic membranes (24). However, the substitution approach is completely different from that used for cellulose acetate. For Hemophan®, only a small percentage (less than 5%) of the hydroxyl groups is actually replaced. However, the tertiary amine 228
Structural And Performance Characteristics Of Hemodialysis Membranes - Clark
replacement group is bulky and effectively shields a significantly greater percentage of hydroxyl groups by a steric mechanism. The attenuation in the degree of complement activation and leukopenia with Hemophan® dialyzers is similar to that observed with cellulose acetate dialyzers. This same approach of providing a low degree of hydroxyl substitution with a relatively bulky moiety is employed for synthetically modified cellulose (SMC), a more recently developed membrane for which the substitution group is a benzyl moiety (25).
4.3. Synthetic Dialyzers Synthetic membranes were developed essentially in response to concerns related to the narrow scope of solute removal and the pronounced complement activation associated with unmodified cellulosic dialyzers. The AN69® membrane, a copolymer of acrylonitrile and an anionic sulfonate group, was first employed in fiat sheet form in a closed-loop dialysate system in the early 1970's (2). Since that time, a number of other synthetic membranes have been developed, including polysulfone (3), polyamide (26), PMMA (27), polyethersulfone (28), and polyarylethersulfone/polyamide (29). Largely related to the interest in hemofiltration (HF) as an ESRD therapy in the late 1970's and early in the following decade, along with the inability to use low-flux unmodified cellulosic dialyzers for this therapy, these membranes were initially formulated with high water permeability (3, 30). The large mean pore size and thick wall structure of these membranes allowed the high ultrafiltration rates necessary in HF, which are typically an order of magnitude higher than those employed in hemodialysis, to be achieved at relatively low transmembrane pressures. However, with the waning of interest in HF as a chronic dialysis therapy in the late 1970's and early in the following decade, dialyzers with these highly permeable membranes were used subsequently in the diffusive mode as high-flux dialyzers. This latter mode continues to be the most common application of these membranes, although they are increasingly being employed for chronic hemodiafiltration (HDF) also (31). Synthetic dialyzers with relatively low water permeability (32) also are now utilized widely in certain markets. An obvious difference between synthetic and cellulosic membranes is chemical composition. As opposed to naturally occurring cellulose, synthetic membranes are manufactured polymers that are classified as thermoplastics. In fact, for most of the synthetic membranes, the hemodialysis market represents only a small fraction of their entire industrial utilization. As noted above, another feature differentiating cellulosic and synthetic membranes is wall thickness. Synthetic membranes have wall thickness values of at least 20 ~tm and may be structurally symmetric (e.g., AN69®, pPMMA) or asymmetric (e.g., polysulfone, polyamide, polyethersulfone, polyamide/polyarylethersulfone). In 229
Structural And Performance Characteristics O f Hemodialysis Membranes - Clark
the latter category, a very thin "skin" (approximately 1 ~tm) contacting the blood compartment lumen acts primarily as the membrane's separative element with regard to solute removal. The structure of the remaining wall thickness ("stroma"), which determines a synthetic membrane's thermal, chemical, and mechanical properties, varies considerably among the various synthetic membranes (14, 33). For example, the stroma has a relatively homogeneous, sponge-like consistency in the Fresenius polysulfone membrane. On the other hand, the Gambro Polyflux membrane, which is actually a blend of polyamide and polyarylethersulfone, has three distinct layers. Similar to the Fresenius polysulfone membrane, one layer is a thin blood-contacting inner lumen, composed of polyarylethersulfone enriched with polyvinylpyrollidone (PVP), while the outer surface is composed of relatively PVP-free polyamide. (See below for discussion about PVP.) Interspersed between these two layers is a polyarylethersulfone stroma (33). Finally, the DiaPES polyethersulfone membrane, developed by Membrana GmbH, contains both inner and outer skin layers surrounding a sponge-like stroma (34). For this membrane, the average pore radius of the inner and outer skin layers are approximately 5 and 10 nm, respectively. Takeyama and Sakai (12; Figure 2) have suggested an effective mean pore radius between 5 and 8 nm achieves the appropriate balance between 132M removal and albumin loss. Many of the synthetic polymers used in the manufacturing of the above membranes are hydrophobic and require the addition of a hydrophilic agent (PVP) to avoid excessive protein adsorption upon blood exposure. During membrane preparation, water-soluble PVP is blended into the hydrophobic base polymer through the formation of hydrogen bonds (26). This hydrogen bonding provides miscibility for the hydrophilic-hydrophobic mixture and prevents leaching of PVP during blood contact. In addition to imparting hydrophilicity, PVP also influences membrane pore size distribution through its effects on pore surface tension and its content may vary considerably across a synthetic membrane's wall thickness. 5.
SUMMARY
A general overview of membrane-related determinants of dialyzer performance has been provided, along with a specific discussion of some of the more commonly used membranes and dialyzers in the HD market. This should assist the reader in understanding the differences that exist between membranes with respect to clinical parameters, such as biocompatility and solute removal. REFERENCES
[1]
WR Clark, Hemodialyzer membranes and configurations: A historical perspective, Sem Dial 13 (2000):309.
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• Structural And Performance Characteristics Of Hemodialysis Membranes
t21 [31 [4]
[5] [6]
[7]
[8] [9]
[10] [11] [12] [13] [14] [15]
[16]
[17]
[18] [19]
[20] [21]
-
Clark
J Funck-Bretano, A Sausse, NK Man, A Granger, M Rondon-Nucete, J Zingraff, P Jungers, A new hemodialysis treatment associating a membrane highly permeable to middle molecules with a closed circuit dialysate system, Proc EDTA 19 (1972):55. E Streicher, H Schneider, The development of a polysulfone membrane: A new perspective in dialysis, Contr Nephro146 (1985): 1. FK Port, SM Orzol, PJ Held, RA Wolfe, Trends in treatment and survival for hemodialysis patients in the United States, Am J Kidney Dis 32/Suppl. 4 (1998):$34. W Bloembergen, RM Hakim, D Stannard, PJ Held, RA Wolfe, L Agodoa, FK Port, Relationship of dialysis membrane to cause-specific mortality, Am J Kidney Dis 33 (1999):1. Y Koda, S Nishi, S Miyazaki, S Haginoshita, T Sakurabayashi, M Suzuki, S Sakai, Y Yuasa, Y Hirasawa, T Nishi, Switch from conventional to high-flux membrane reduces the risk of carpal tunnel syndrome and mortality of hemodialysis patients, Kidney Int 52 (1997):1096. H Strathmann, H Gohl, Membranes for blood purification: State of the art and new developments, Contrib Nephrol 78 (1990): 119. CK Colton, EG Lowrie: Hemodialysis: Physical principles and technical considerations; in BM Brenner, FC Rector (eds): The Kidney, 2nd edition. Philadelphia, Saunders, 1981, pp. 2425-2489. RB Bird, WE Stewart, EN Lightfoot: Velocity distributions in laminar flow; in RB Bird, WE Stewart, EN Lightfoot (eds): Transport Phenomena, 1st edition. New York, John Wiley and Sons, 1960, pp. 34-70. WR Clark, Quantitative characterization of hemodialyzer solute and water transport, Sem Dial 14 (2001):32. G Lonnemann, Chronic inflammation in hemodialysis: The role of contaminated dialysate, Blood Purif 18 (2000):214. T Takeyama, Y Sakai, Polymethylmethacrylate: One biomaterial for a series of membrane, Contrib Nephrol 125 (1998):9. WR Clark, C Ronco, Determinants of hemodialyser performance and the potential effect on clinical outcome, Nephrol Dial Transplant 16/Suppl. 5 (2001):56. H Sugaya, Y Sakai, Polymethylmethacrylate: From polymer to dialyzer, Contrib Nephrol 125 (1998):1. WR Clark, RJ Hamburger, MJ Lysaght, Effect of membrane composition and structure on performance and biocompatibility in hemodialysis, Kidney Int 56 (1999):2005. C Ronco, M Scabardi, M Goldoni, A Brendolan, C Crepaldi, G LaGreca, Impact of spacing filaments external to hollow fibers on dialysate flow distribution and dialyzer performance, Int J Artif Organs 20 (1997):261. C Ronco, A Brendolan, C Crepaldi, M Rodighiero, M Scabardi, PM Ghezzi, Blood and dialysate flow distributions in hollow fiber hemodialyzers analyzed by computerized helical scanning technique, J Am Soc Nephrol 13 (2002):$53. MJ Lysaght, Evolution ofhemodialysis membranes, Contrib Nephrol 113 (1995): 1. MJ Lysaght, Hemodialysis membranes in transition, Contrib Nephrol 61 (1988): 1. KK Jindal, J McDougall, B Woods, L Nowakowski, MB Goldstein, A study of the basic principles determining the performance of several high-flux dialyzers, Am J Kidney Dis 14 (1989):507. WR Clark, WL Macias, BA Molitoris, NHL Wang, Plasma protein adsorption to highly permeable hemodialysis membranes, Kidney Int 48 (1995):481.
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Structural And Performance Characteristics Of Hemodialysis Membranes - Clark
[22]
[23] [24] [25] [26] [27]
[28]
[29]
[3o1 [31]
[32] [331 [34]
M Grooteman, M Nube, J van Limbeek, A van Houte, M Daha, J van Geelen, Biocompatibility and performance of a modified cellulosic and a synthetic high-flux dialyzer, ASAIO J 41 (1995):215. N Hoenich, C Woffindin, P Cox, M Goldfinch, S Roberts, Clinical characterization of Dicea a new cellulose membrane for haemodialysis, Clin Nephro148 (1997):253. R Schaefer, WHorl, K Kokot, A Heidland, Enhanced biocompatibility with a new cellulosic membrane: Cuprophan vs Hemophan, Blood Purif 5 (1987):262. WR Clark, JH Shinaberger, Clinical evaluation of a new high-efficiency hemodialyzer: Polysynthane (PSNrM), ASAIO J 46 (2000):288. H Gohl, R Buck, H Strathmann, Basic features of polyamide membranes, Contrib Nephrol 96 (1992): 1. M Bonomini, B Fiederling, T Bucciarelli, V Manfrini, C Di Ilio, A Albertazzi, A new polymethylmethacrylate membrane for hemodialysis, Int J Artif Organs 19 (1996):232. BL Jaber, JA Gonski, M Cendoroglo, VS Balakrishnan, P Razeghi, C Dinarello, BJG Pereira, New polyethersulfone dialyzers attenuate passage of cytokine-inducing substances from pseudomonas aeruginosa contaminated dialysate, Blood Purif 16 (1998):210. NA Hoenich, S Stamp, Clinical performance of a new high-flux synthetic membrane, Am J Kidney Dis 36 (2000):345. A Rockel, J Hertel, P Fiegel, S Abdelhamid, N Panitz, D Walb, Permeability and secondary membrane formation of a high flux polysulfone hemofilter, Kidney Int 30 (1986):429. I Ledebo, Principles and practice of hemofiltration and hemodiafiltration, Artif Organs 22 (1998):20. R Ward, A Buscaroli, B Schmidt, S Stefoni, H Gurland, H Klinkmann, A comparison of dialysers with low-flux membranes: Significant differences in spite of many similarities, Nephrol Dial Transplant 12 (1997):965. GJ Mishkin, What clinically important advances in understanding and improving dialyzer function have occurred recently?, Sem Dial 14 (2001): 170. BC Li, CK Poh, Z Huang, Z Liao, D Gao, WR Clark, Cytokine-inducing substance backtransfer for a new high-flux polyethersulfone dialyzer (abs), J Am Soc Nephrol 12 (2001):268A.
232
New Insights into Membrane Science and Technology: Polymeric and Biofunctional Membranes D. Bhattacharyya and D.A. Butterfield (Editors) © 2003 Elsevier Science B.V. All rights reserved.
Chapter 11
Biofunctional membranes: site-specifically immobilized enzyme arrays D. A. Butterfield ~*, D. Bhattacharyya ~c aDepartment of Chemistry, bDepartment of Chemical and Materials Engineering, and c Center of Membrane Sciences, University of Kentucky, Lexington, KY 40506-0059 U.S.A. *Corresponding Author: Professor D. Allan Butterfield, Department of Chemistry and Center of Membrane Sciences, University of Kentucky, Lexington, KY 40506-0059 e-mail: [email protected] 1.
INTRODUCTION
Biofimctional membranes are entities in which a biomolecule, collection of biomolecules or cells are immobilized onto polymeric matrices cast in the form of porous membranes [1, 2]. Biofunctional membranes are used in catalysis [membrane-based enzyme bioreactors], separations [affinity membranes], analysis [biosensors; metal ion-specific separations], and artificial organs [1, 2]. This chapter deals with the production and analysis of catalytic biofimctional membranes in which enzymes are site specifically immobilized to the membrane in processes that combine physical chemical, chemical engineering, and molecular biology methods. The stability of enzymes, relative to those in solution, is enhanced by membrane immobilization [1, 3-6]. However, the activity of immobilized enzymes on polymeric membranes most often decreases, a problematic result due to random immobilization of enzymes in which the active site of the immobilized enzyme is oriented in different directions. Steric hindrance of the active site from substrate accessibility, multiple-point binding, or denaturation of the enzyme are some of the consequences of random immobilization of enzymes and causes the low activity [6-11] (Figure 1A). Enzymes are directly attached onto the membrane or to a spacer arm, usually via the e-amino group of lysine residues on the protein. However, given the often large number of lysine residues on the surface of the enzyme, this random immobilization strategy results in different orientations of the enzyme and to denaturation of enzyme catalytic sites as a consequence of protein-surface interactions. The principles of molecular biology often can lead to obviation of these problems associated with random immobilization of enzymes [8]. Site-specific
233
BiofunctionalMembranes:Site-SpecificallyImmobilizedEnzymeArrays-Butterfield&Bhattacharyya immobilization of enzymes results in a two-dimensional array of proteins with their active sites all oriented away from the membrane surface [8]. In our laboratories, ordered arrays of enzymes on membrane surfaces have been produced using the following molecular biology methods: (a) Gene fusion to incorporate a peptide affinity tag at the N- or C-terminus of the enzyme; the enzymes are then attached from this affinity tag to anti-tag antibodies on membranes; (b) Posttranslational modification to incorporate a single biotin moiety on enzymes; the enzymes can be attached through a (strept)avidin bridge; (c) Site-directed mutagenesis to introduce unique cysteines to enzymes; the enzymes are attached on thiol-reactive surfaces through the sulfhydryl group on the side chain of the introduced cysteine. In the latter case the SH group is introduced to the enzyme on the opposite side of the protein from the active site. In all these methods, the active sites of the immobilized enzymes face away from the polymeric surface and, as we demonstrated, a consequent higher activity was retained (reviewed in [8]).
~_ .-V} ....
/i~
. . . . . . . . . .
\_
_
_/_.
i:!:i:!:i:!:i:i:i:i:i:i:i:!:i:i:i:i:i:i:i:!:!:i:i:i:i:!:!!!:!:!l Figure 1A. Random immobilization of proteins. Indentation indicates binding/active site of the protein.
~
,
Antibody, ~ '~/~ Antib°dy4.[ Pr°teinA I
Figure lB. Protein A and anti-FLAG monoclonal antibody mediated site-specific immobilization of FLAG -tagged proteins. Note that the active site of all enzymes faces away from the polymeric membrane surface and towards the solution. 2. SITE-SPECIFIC IMMOBILIZATION MOLECULAR BIOLOGICAL METHODS
OF
ENZYMES
USING
Production of a highly ordered, two-dimensional arrays of enzyme on membranes is desired (Figure IB). As noted above, this arrangement leads to
234
Biofunctional Membranes: Site-SpecificallyImmobilized Enzyme Arrays -Butterfield & Bhattacharyya
orienting the active site away from the polymeric surface, thereby giving the substrate nearly full accessibility. Moreover, because the enzyme is immobilized, greater resistance to denaturants, pH, and temperature is produced [1]. In cell membranes, proteins are oriented, and one goal of biofunctional membranes is to mimic as closely as possible the real systems [ 12]. A number of methods to achieve oriented immobilization of enzymes are known [13]. Biotin has an affinity for avidin that is remarkably strong: the KD of the avidin-biotin complex is approximately 1015. Molecular biology can be used to accomplish biotin addition to an enzyme in a site-specific fashion. For example, 13-galactosidase ([3-Gal) can be biotinylated at a single site following expression of a plasmid in E. coli that encodes a fusion protein of 13-Gal and a polypeptide tag at its N-terminus. Expression of this fusion protein in E. coli leads to one lysine residue on the polypeptide tag being specifically biotinylated by the bacteria-resident enzyme, biotin ligase, during the posttranslational modification process. Immobilization to poly(ether) sulfone membranes onto which avidin was already immobilized yielded a two-fold enhancement of enzyme activity relative to that of fully biotinylated enzyme, and a nearly 20fold enhancement of activity was found compared to randomly immobilized 13galactosidase [ 14]. Gene fusion methods also were used to effect site-specifically immobilization of alkaline phosphatase (AP) [15]. A fusion protein of AP and FLAG (Asp-Tyr-Lys-Asp-Asp-Asp-Asp-Lys) was reacted with a monoclonal antibody directed against FLAG. Protein A, which has a high binding towards and specificity for monoclonal IgG anti-FLAG antibody, was previously immobilized on the membrane. Site-specific orientation of AP resulted in higher catalytic activity relative to the randomly immobilized enzyme, (over 85% of the activity of the soluble enzyme) [15]. A different modular biological method was used for subtilisin. Site directed mutagenesis was employed to put a cysteine into subtilisin [3, 4], a protein with no cysteines. Subtilisin has serine residues on the distal side of the protein away from the active site serine of this protease. Site-directed mutagenesis introduced a cysteine, one at a time, in these remote serines. This is a conservative mutation for serine, since the OH functionality of serine is replaced by SH. One cysteine was introduced into the protease, and kinetic analysis of the mutated enzyme in solution showed that no significant difference in activity. Further, EPR studies (see below) indicated that there was no change in mobility of a spin label bound in the active site, i.e., no change in active site structure. Subsequent site-specific immobilization of enzymes modified by sitedirected mutagenesis employs the unique cysteine located away from the active site. The SH group is coupled to a thiol-reactive membrane support (Figure 2). A significant increase in relative activity (up to 83% compared to that of the
235
Biofunctional Membranes: Site-SpecificallyImmobilizedEnzymeArrays-Butterfield & Bhattacharyya
soluble enzyme) was observed for site-specifically immobilized subtilisin on PVC-silica membranes [ 16].
PCR
Expression
SH
Immobilization S I-B-BB~--m-_ •
FigtLre2. S i t e - ~ c immob'tlizationof enzymesusing site-directedmutagenesis. To a protein with no cysteines, this amino acid is added in lieu of serine at a site distal from the active site of the enzyme. Subsequentreaction of the enzymewith the surfacecontaininga SH reactivemoeityleads to site-specific immobilizationof the enzyme. See text. Recharging catalytic biofunctional membranes is easily accomplished using gene fusion methods. The enzyme and calmodulin, a Ca2+-binding protein with high affinity for phenothiazine in the presence of Ca 2+, are formed as a fusion protein complex by recombinant DNA methods [17]. The complex is added to the membrane support in the presence of C a-2+ , leading to site-specific immobilization. When the enzyme is spent, EGTA, a specific Ca 2+ chelator is added. This results in a rapid and complete separation of calmodulin and attached enzyme from phenothiazine. To recharge the system, a new calmodulin-enzyme fusion protein is added in the presence of Ca 2+. This approach is appropriate when the membrane cost is high since expensive supports can be reused. A higher activity per mg bound enzyme relative to randomly immobilized enzyme is found using site-specifically immobilized enzyme arrays [16]. This probably is a consequence of reduced steric hindrance of the active site in enzyme arrays by adjacent enzyme molecules and by the membrane surface since each enzyme molecule is oriented in the same manner with the active site facing away from the polymeric membrane surface (Figure 1B). Increased enzyme activity, higher loading, and an active-site structure similar to that of the soluble enzyme results from site-specifically immobilized enzymes. That oriented enzymes with the active site directed away from the polymer surface are predicted to demonstrate greater enzyme activity was confirmed using the trans-esterification reaction catalyzed by subtilisin in an organic solvent. The rate of the reaction of 1-butanol with CBZ-alanine-p-
236
Biofunctional Membranes: Site-Specifically Immobilized Enzyme Arrays -Butterfield & Bhattacharyya
nitrophenyl ester in water-saturated hexane was determined for randomly immobilized and site-specifically immobilized subtilisin that had been mutated to introduce a cysteine in a location distal from the active site. This was a stringent test of our thesis, since the protease activity of subtilisin is greater than its esterase activity, and since the reaction occurs in organic media, a lower rate would be expected. However, consistent with our hypothesis, a two-fold enhancement of the rate was found for subtilisin that had been subjected to sitedirected mutagenesis and site-specifically immobilized onto PVC-silica membranes [3, 16]. 3.
ANALYSIS OF ACTIVE SITE ORIENTATION
To evaluate the efficiency of the site-specifically immobilized enzymes catalytic activity often is used. However, if optical methods of analysis are used, problems can arise due to light scattering by membrane surfaces. We developed a novel approach to measuring enzyme activity of randomly and sitespecifically immobilized enzymes on membranes. Electron paramagnetic resonance (EPR), which is not affected by light scattering, was shown to be highly effective in measuring enzyme activity, comparable to traditional methods. This magnetic resonance analysis is based on determining the difference in intensity of an active site-specific spin label before and after reaction with the immobilized enzyme. The difference in intensity is hypothesized to result from the accessibility of active site of the enzyme to spin label molecules. We earlier showed that only non-denatured and accessible active sites incorporated active-sites spin labels [7]. Hydrophilic and hydrophobic membranes, bacterial cellulose and modified poly(ether)sulfone (MPS) membranes, respectively, were used in both random and site-specific immobilization techniques. Subtilisin and organophosphorus hydrolase (OPH) were used to generalize the findings to enzymes of different types. While subtilisin is a serine protease [18], OPH has two divalent metal ions in its active site [ 19]. Knowing the amount of active enzyme site-sPecifically immobilized on the membrane and whether this amount is correlated to enzyme activity would provide a means of evaluating different methods of enzyme immobilization and the effects of the membrane on the activity of the enzyme. As noted above, these principles were evaluated using two markedly different membranes, one hydrophobic [modified poly(ether)sulfone] and one hydrophilic (bacterial cellulose) and two different types of enzymes, a serine protease and a metalcentered hydrolase. Further, site-specific immobilization of both enzymes was compared to random immobilization.
237
Biofunctional Membranes: Site-Specifically Immobilized Enzyme Arrays -Butterfieid & Bhattacharyya
4.
PRINCIPLES OF THE SPIN LABEL TITRATION METHOD.
Only accessible active sites in non-denatured enzymes will incorporate the spin label specific for each type of enzyme [7]. Analysis of the EPR spectrum of a known concentration of spin label before and after reaction with the enzyme gives the amount of spin label incorporated [20]. Companion to a standard takes into account possible intensity variations resulting from sample placement differences in the EPR resonant cavity [20]. The results (Table 1) demonstrate three important findings. (a) The spin label titration method gives excellent agreement with activity measurements. Thus, this method, applicable to a wide variety of enzymes, provides a means of measuring membrane-bound enzyme activity without measuring enzyme activity. This is a solution to the potential problem of light scattering in activity measurements of membrane immobilized enzymes. (b) Site specific immobilization gives extraordinarily higher activities (as assessed by both EPR and kinetics) than does random immobilization. This is undoubtedly due to all copies of the immobilized enzyme oriented such that their active sites are directed away from the membrane surface and to the higher enzyme loading possible with site-specific immobilization [ 1-4,21 ]. This argues strongly for the power of molecular biology to enhance catalytic biofunctional membrane performance. (c) The enzymes immobilized to hydrophilic membranes give much higher activity than that bound to hydrophobic MPS membranes. This is likely due to decreased enzyme-polymer interactions in the more polar membrane [3]. Table 1. Comparison of Spin Label Titration (SLT) and the Activity Method for Determining Active Immobilized Enzyme on MPS and Bacterial Cellulose Membranes a. % Active Enzyme OPHImmobilization Membrane Method Subtilisin Subtilisin OPH FLAG -FLAG Technique
11.7+2.7 9.4+1.9 SLT 8.0+5.2 Activity 10.6+4.3 ~34.9+1.5 31.5+4.0 Random BC SLT 37.0 Activity 27.4 51.0+1.6 28.5+1.7 MPS SLT Site Specific 49.0+7.2 28.1+6.8 Activity 84.3+1.2 82.5+2.6 BC SLT Site Specific m 89.0+9.2 80.6+9.1 Activity a. The results are given in percentage of activity of the respective enzyme in homogenous solution. N = 2-4 for each measurement.
Random
MPS
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Biofunctional Membranes: Site-Specifically Immobilized Enzyme Arrays -Butterfield & Bhattacharyya
5.
CONCLUSION
Multidisciplinary approaches were combined to gain insights into sitespecifically immobilized biofunctional membranes. Greater understanding and application of catalytic biofunctional membrane enzyme arrays for catalysis, analysis, and separation is envisaged. Future studies may take advantage of directed evolution to produce site-selectively immobilized proteins of even more enhanced activity. EPR has predictive power for biofunctional membrane immobilization and functional effectiveness. Moreover, future use of functionalized solvent-resistant membranes in non-aqueous solvents conceivably could yield enhanced biocatalytic reaction selectivity. Such studies are ongoing in our Center of Membrane Sciences. 6.
ACKNOWLEDGEMENTS This work was supported in part by a grant from DoD (DAAG55-98-1-
0003). REFERENCES
[1] [2] [3]
[4] [5] [6] [7]
[8] [9] [101 [11] [12] [13] [14]
[15]
D. A. Butterfield(editor), Biofunctional Membranes; Plenum Press: New York, 1996. G. F. Bickerstaff (editor), Immobilization of Enzymes and Cells; Humana Press: Totowa, New Jersey, 1997. S. Viswanath, J. Wang, L. G. Bachas, D. A. Butterfield, D. Bhattacharyya, Biotechnol. Bioeng. 60 (1998) 608-616. J. Wang, D. Bhattacharyya, L. G. Bachas, Biomacromol. 2 (2001) 700-705. D. S. Clark, TIBTECH 12 (1994) 439-443. A. Bhardwaj, J. Lee, K. Glauner, S. Ganapathi, D. Bhattacharyya, D. A. Butterfield, J. Membr. Sci. 119 (1996) 241-252. D. A. Butterfield, J. Lee, S. Ganapathi, D. Bhattacharyya, J. Membr. Sci. 91 (1994) 47-64. D. A. Butterfield, D. Bhattacharyya, S. Daunert, L. G. Bachas, J. Membr. Sci. 181 (2001) 29-37. C. C. Tsai, Y. Chang, H. W. Sung, J. C. Hsu, C. N. Chen, Biomaterials 22 (2001) 523533. R. Vankova, A. Gaudinova, H. Sussenbekova, P. Dobrev, M. Strnad, J. Holik, J. Lenfeld, J. Chromatogr. A 811 (1998) 77-84. V. V. Shmanai, T. A. Nikolayeva, L. G. Vinokurova, A. A. Litoshka, BMC Biotechnol. 1 (2001) 4. J. Liu, J. Wang, L. G. Bachas, D. Bhattacharyya, Biotechnol. Prog. 17 (2001) 866871. R.F. Taylor (Ed.) (1991) Protein Immobilization: Fundamentals and Applications, Marcel Dekker, Inc., New York. S. Vishwanath, D. Bhattacharyya, W. Huang, L.G. Bachas, J. Membr. Sci. 108, (1995) 1. S.K. Vishwanath, C.R. Watson, W.Huang, L.G. Bachas, D. Bhattacharyya, J. Chem. Technol. Biotechnol. 68(1997) 294.
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Biofunctional Membranes: Site-Specifically Immobilized Enzyme Arrays -Butterfield & Bhattacharyya
[16] [17] [18] [19] [20] [21]
S. Vishwanathl J. Wang, L.G. Bachas, D.A. Butterfield, D. Bhattacharyya, Biotech. Bioeng. 60 (1998) 608. V. Schauer-Vukasinovic, S. Daunert, Biotechnol. Prog. 15(1999) 513. G. L. Gilliland, D. T. Gallagher, P. Alexander, P. Bryan, Adv. Exp. Med. Biol. 379 (1996) 159-169. J. L. Vanhooke, M. M. Benning, F. M. Raushel, H. M. Holden, Biochemistry 35 (1996) 6020-6025. D.A. Butterfield, J. Colvin, J. Liu, J. Wang, L. Bachas, and D. Bhattacharrya, Anal. Chim. Aeta (2002) in press. S. Ganapathi, D. A. Butterfield, D. Bhattaeharyya, Biotechnol. Prog. 14 (1998) 865.
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New Insights into Membrane Science and Technology: Polymeric and Biofunctional Membranes D. Bhattacharyya and D.A. Butterfield (Editors) © 2003 Elsevier Science B.V. All rights reserved.
Chapter 12
Biocatalytic membrane reactor with continuous removal of organic acids by electrodialysis H. C. Ferraz, T. L. M. Alves, C. P. Borges* Chemical Engineering Program- COPPE, Federal University of Rio de Janeiro, P.O.Box: 68502, 21945-970, Rio de Janeiro / R J - Brazil *Corresponding author, e-mail: [email protected] 1.
INTRODUCTION
In the last decades, a constant increase in the production of chemical compounds through biotechnological processes has been observed. In this sense, scientific and technological advances in microbiology, as well as in bioreactors engineering were fundamental. The main advantages to apply biosynthesis in technical scale are mild operation conditions, low energy consumption, high specificity minimizing by-products and reduced environmental impact. However, the majority of bioprocesses are still operated in batch or fed-batch regimes that, in general, lead to low productivity and higher operational costs. Biocatalysts such as enzymes and cells exhibit specific and effective catalytic action well beyond the performance of synthetic catalysts. The difficult and expensive recovery and reuse of enzymes or cellular microorganisms have motivated the development of immobilization techniques. In these techniques the biocatalyst is confined in a well-defined space, keeping their catalytic properties, and making possible the repeated and continuous use. Furthermore, immobilization simplifies the downstream operations, leads to more compact plants and enhances the stability of the biomaterial. The immobilization of the biocatalyst also facilitates coupling of separation processes to the reactor allowing continuous removal of products and, consequently, a better performance. Usually, product accumulation inhibits the bioactivity or even causes irreversible inactivation of enzymes and microorganism. Among the separation processes, the membrane technology seems to be one of the best solutions to be coupled in bioreactors. The separation is carried out continuously and in conditions similar to the reaction medium. The distinction of species is based on size or mobility and, frequently, at membrane surface. These characteristics, the variety of membrane materials and their large available surface per unit volume also permit new possibilities for
241
Biocatalytic Membrane Reactor With Continuous Removal Of Organic Acids By Electrodialysis - Borges
immobilization procedures. The biocatalyst can be immobilized by binding at the membrane surface or it can be confined in a defined region by size exclusion. The membrane with the biocatalyst is named as catalytic membrane reactor and presents several advantages with respect to conventional ones, such as combination of selective mass transfer across membrane with biochemical reactions. In the present chapter, biocatalytic membrane reactor for enzymatic sugar conversion to carboxylic acid is described. In particular, the production of gluconic acid is examined. In the investigated biosynthesis the production of gluconic acid occurs from glucose oxidation, which needs a simultaneous reaction to regenerate the enzyme cofactor. This parallel reaction reduces fructose to sorbitol. Hence, using a glucose/fructose mixture as substrate it is produced gluconic acid and sorbitol in equimolar ration. An electrodialysis unit is also coupled to the reaction system for continuous removal of the acid produced, which avoids pH reduction and enzyme inactivation. A brief description of the gluconic acid biosynthesis using permeabilized cells will be also given. 2.
IMMOBILIZATION OF BIOCATALYSTS
Biotransformations constitute an appealing altemative to conventional chemical process for production of some classes of chemicals, such as antibodies, chiral molecules, organic acids, etc. Types of reactors used for this purpose, called bioreactors, include immobilized cell, two-phase contactors and membrane bioreactors. A large number of processes using immobilized biocatalysts have been investigated due to innumerous advantages of this type of system [1], [2]. Comparing to the use of free cells, immobilization facilitates the operation in continuous mode, with all its inherent advantages: better control, better product uniformity, removal of potentially toxic metabolites, etc. Cells immobilization can, also, stabilize the biocatalyst over extended period. Desirable features in a support are: absence of toxicity, high capacity of retention, stability in the desired range of temperature and pH, low cost and reusable. Main methods of immobilization are occlusion, adsorption, flocculation, covalent binding and encapsulation. The confinement by size exclusion using membranes is an emerging technology with few examples reported [2]. Adsorption in ion exchange resin can raise problems of cells leakage, mainly during fermentation. Polyurethane foams are versatile, inert and with low cost. They have an open structure allowing a more effective diffusion. Covalent binding to the active support generally results in death of cells due to the toxicity of some of the agents employed. However, it does not present diffusion limitation and it is less subject to the cells loss. Using flocculation, as 242
Biocatalytic Membrane Reactor With Continuous Removal Of Organic Acids By Electrodialysis - Borges
no solid support is required, the density of cells is high. However, the flakes do not present much resistance. Cells occlusion in gel beads is the most widely used immobilization method. In this case, gel is formed in presence of the cells. Frequently, a crosslinking agent like glutaraldehyde is also used, which improves cells retention and stabilizes the gel. The use of synthetic polymers like polyacrilamide can result in cells death due to high toxicity. Natural polymers, such as alginate, kcarrageenan, agar, gelatin and chitosan, generally preserve cells viability after immobilization. Alginate is a polysaccharide obtained from kelps, a copolymer of 13-Dmanuronate and c~-L-guluronate. Cells immobilization is achieved by ionotropic gelation in presence of ions Ca 2+ or other cations such as A13+, Ba z+ and Cu 2+. An aqueous solution containing alginate and cells form gel through crosslinking. A limitation of the systems is its ready especially in presence of chelating gents, as phosphate, citrate and EDTA. Carrageenan is a 13-Dgalactose-4-sulphate and 3,6-anidro-D-galactose copolymer. It forms rigid gels by cooling. It is very used in food industries because of its absence of toxicity. An interesting application of the immobilization technique is the construction of biosensors. Biosensors are instruments that combine a biologically active material with a transducer, converting the answer to a specific chemical species into a measurable electric sign. They present important advantages in relation to other analytical methods: high specificity, fast reply time and possibility of online measurement [3]. In two-phase bioreactors, two liquid phases (aqueous/aqueous, aqueous/organic or organic/organic) are in direct contact. The most common is the aqueous/organic system. Inconvenience of this process is the formation of emulsion due to presence of biological surface-active materials. Membrane bioreactors are an alternative to direct two-phase contactors. A microporous [4] or a dense [5] membrane constitutes a fixed interface between organic and aqueous phases, allowing contact without emulsification. The biocatalyst is usually present in two forms: immobilized or confined inside the membrane pores. In the former immobilization of the biocatalyst can be achieved by chemical bounding, adsorption or electrostatic attraction at membrane surface. In confinement, the membrane is used a selective barrier to retain the soluble biocatalyst [2], [6]. Hollow fiber membranes have been used in process like reverse osmosis, micro and ultrafiltration. In biotechnology, growing cells in hollow fiber membrane bioreactors (HFMB) has also been done for more than 30 years. More recently, HFMB are used for production of many proteins, cells and bioartificial organs [7], [8]. Depending on the immobilization type, membrane configuration, and fluid dynamics, the following main catalytic membrane reactor configurations might be considered [2]: 243
Biocatalytic Membrane Reactor With Continuous Removal Of Organic Acids By Electrodialysis - Borges
-
Enzyme membrane reactors and fermentors, where the biocatalyst are continuously flushed along the membranes;
-
Membrane reactors with confined biocatalyst, where the biocatalyst is confined within the membrane module (in the lumen or in the shell side) or entrapped within the membranes pores;
-
-
Membrane reactor with gelified biocatalyst, where the immobilization occurs in a proteic gel layer, dynamically or statically formed; Membrane reactor where the biocatalyst is chemically bound to the membrane surface.
Compared to the conventional methods, use of hollow fiber present advantages such as [9]: -
High volumetric productivity, as consequence of the great packing density;
-
Simultaneous separation of products;
-
Prevents problems of cells loss;
-
It is easily reutilized;
-
Cells growth and immobilization can be carried out in a single step once cells can be cultivated inside the bioreactor;
-
Lower sterilization costs, once many fibers can be chemically sterilized.
However, some limitations associates to the use of this type of bioreactor are: -
Solutions must be free of particulate material;
-
Restricted use of organic solvent, depending on the type of fiber;
-
Problems with fiber rupture, mainly when there is extreme growth or gas evolution.
Many processes using cells of Saccharomyces cerevisiae, Lactobacillus delbreuckii, Lactobacillus bulgaricus, Streptococcus thermophilus, and several other microorganisms have been carried out successfully in this type of system [9]. Recent applications of biocatalytic membrane reactors are described by Drioli and Giomo [2], including production of chirals and use as artificial organs. B I O R E A C T O R COUPLED TO M E M B R A N E P R O C E S S E S P R O D U C T I O N AND R E C O V E R Y OF CARBOXYLIC ACIDS 3.
FOR
The combination of a bioreactor with separation processes is particularly necessary when the product needs further processing for concentration and
244
Biocatalytic Membrane Reactor With Continuous Removal Of Organic Acids By Electrodialysis - Borges
purification. The usual low product concentration involved in bioprocesses may turn the product recovery rather complicated. The downstream processing may include several unit operations and generally represents a large part of the overall capital investment in a biotechnological process. In some cases it can be determinant to its economical viability. Many studies with bioprocesses coupled to separation systems have point out that membrane technology is one of the best approach forms to improve economical and technical competitiveness. When carboxylic acids are produced, besides the design of an appropriated downstream processing, it is important to maintain its concentration as low as possible in the reaction environment to avoid pH reduction with consequent biocatalyst inactivation. The most common membrane separation processes used for production of carboxylic acid will be discussed in this section. Carboxylic acids, such as citric, malic, lactic, tartaric and gluconic acids, are largely used as additives in the food industry. Citric and malic acids are used to adjust the acid flavor in soft drinks, fruit juices and wine. In a certain concentration they inhibit the development of undesirable flavors and color deterioration. Lactic acid and calcium lactate are used to adjust acidity in wines and in the preparation of dried milk powers, respectively. A great number of applications indicates a huge market and indicates the necessity to develop processes to produce them with high efficiency and at low cost. The usual production of carboxylic acid is carried out in batch reactors, using substrates that, in general, include glucose, sucrose, beet molasses or cane molasses. Concentration and purification of the acid is effectuated at the end of the fermentation. The downstream processing involves a filtration of cells and suspended solids, acidification (to precipitate calcium sulfate), evaporation, metal removal by ion exchange and crystallization. Carboxylic acids can be produced in a continuous way adding the fresh medium continuously to the reactor, while a mixture of cells and fermented medium is withdrawn at same rate. This procedure would avoid the timeconsuming start-up phase of batch reactor, but more efficient production is obtained by confining the biocatalyst in a defined region. Several studies are reported in literature using confining the biocatalyst by immobilization on a substrate surface or entrapped into gel beads [10], [11 ], [12]. Other studies have been investigated a biocatalyst confinement by size exclusion filtration [13], [ 14], [ 15], [ 16], [ 17]. A typical bioreactor coupled to membrane process consists of a fermentor connected to an ultrafiltration (UF) unit. The feed stream with substrate and nutrients is continuously added to the reaction tank. A pump circulates the fermented liquor through the UF unit, where cells debris, macromolecules and solids are retained in the stream by the membrane, returning to the reaction tank. Product and other low molecular weight compounds permeate through the membrane and are removed from the system. The process productivity increases by avoiding or reducing product inhibition, 245
Biocatalytic Membrane Reactor With Continuous Removal Of Organic Acids By Eiectrodialysis - Borges
however the UF should only start when the concentration of substrate and low molecular weight nutrients is decreased to diminish their concentration in the permeate stream. The continuous fed of microorganism leads to high biocatalyst density, which should be controlled by periodical bleed to better operation of the membrane unit. Drioli and Giomo [2] have compared data of several authors using different microorganism in continuous membrane fermentors for lactic acid production. It has been point out that the productivity is, in general, higher than that of batch processes. However, a permeate stream with low concentration of lactic acid decreases the economic advantage of high productivity. Other carboxylic acids were also produced in similar coupled system, where the permeate recovered from UF unit was concentrated by a subsequent nanofiltration unit [ 18]. Others membrane processes, such as electrodialysis, pertraction, ultrafiltration and nanofiltration, have also been investigated as new technology for continuous processing of carboxylic acid. The application of electrodialysis in production of lactic acid was specially to remove metabolites inhibiting the microorganism activity [19]. The same authors point out operation problems with electrodialysis unit, such as strong concentration polarization phenomenon and membrane fouling by cellular material. This later disadvantage can be solved by a previous microfiltration of the medium [20]. An integrated system for production and recovery of lactic acid, consisting of diffusion dialysis (permeation driven force is the concentration gradient) and electrodialysis units operating successively has also been reported [21]. In the dialysis unit the acid permeates to a pH-controlled alkaline solution, neutralizing the fermentation broth. Thus, the sodium lactate is converted into acid in the electrodialysis unit. Besides lactic acid, propionic acid production in bioreactors coupled to electrodialysis unit has also been reported [22]. It was reported the possibility to concentrate large amounts of cells within the system during continuous fermentation and to reduce the cell inhibition effect of organic acid accumulation in the medium. The increased efficiency of membrane processes has considerably facilitated the recovery and purification steps from fermented media. Reverse osmosis (RO) and nanofiltration have been frequently investigated to concentrate low molecular weight compounds, such as carboxylic acids [23]. Lactic acid rejection of RO membranes was found to be strongly dependent on pH, which was related to rejection of dissociated species. Pertraction with liquid membranes, supported liquid membranes and membrane-based solvent extraction has been also used for the separation of organic acids. Organic solvents and specific carriers have been studied for lactic acid extraction [24], [25], [26] showing high fluxes and selectivities. However
246
Biocatalytic Membrane Reactor With Continuous Removal Of Organic Acids By Electrodialysis - Borges
the stability of the membrane still is the major drawback to operate at technical scale. 4.
GLUCONIC ACID AND SORBITOL PRODUCTION
Production of gluconic acid and sorbitol by Zymomonas mobilis has been studied extensively since the pioneer work of Zachariou and Scopes (1988) [27]. As will be seen more detailed latter, this process constitute a possible alternative to the conventional methods of production. Bioreactors such as CSTR [28], fluidized bed [29] and hollow fiber membrane reactor [30] have been used with this purpose. Sorbitol is a polyol (sugar alcohol) found in several food products. It is a bulk sweetener, an excellent humectant and texturizing agent. It is about 60% as sweet as sucrose and with one-third fewer calories. It has a smooth a sweet, cool and pleasant taste and it is non-cariogenic. It is also used by pharmaceutical and cosmetic industries. Sorbitol is very stable and chemically unreactive. Gluconic acid is the polyhydroxycarboxylic acid. It exhibits an excellent chelating power, especially in alkaline and concentrated alkaline solutions. Calcium, iron, copper, aluminum and other heavy metals are firmly chelated in alkaline solution. It is stable at the boiling point even of concentrated alkaline solutions. Gluconic acid and its salts have been used for many applications such as in household cleaner, in industrial cleaner, in inks, paints, dyes, for metal finishing, as paper/textile auxiliaries and in water treatment [31 ]. Sorbitol is produced by catalytic hydrogenation of glucose. Nickel is generally used as catalyst to the reaction, which is carried out at 140°C and 4050 atm, and must be recovered at the end. Purification of sorbitol solution is achieved in two steps: initially through an ion exchange column, to remove gluconic acid and other ions present, and then, a treatment with active coal. Although gluconic acid can be obtained by chemical synthesis from glucose, biosynthesis is more advantageous in this case due to mild operation conditions with minimal by-product formation because of the enzyme specificity, responding for more than 60% of total gluconic acid production. Zachariou and Scopes (1986) identified an enzymatic complex in Zymomonas mobilis capable of producing sorbitol and gluconic acid from glucose and sorbitol simultaneously. Many studies have been conduced since then [28], [29], [32], [33] indicating that this process has a very good perspective to be applied in a larger scale. 5.
CELLS PERMEABILIZATION
Double layers of lipids constitute the structural base for all biological membranes. A typical biological membrane contains 40% of phospholipids and 247
Biocatalytic Membrane Reactor With Continuous Removal Of Organic Acids By Electrodialysis - Borges
glycolipids and the remaining fraction is composed of proteins. Non-covalent interactions between lipids of the layer confer flexibility to the membrane. The plasma membrane represents a barrier to the transport of specific molecules into or out of the cell, besides other important functions. Initially, there was more interest in taking specific cellular components than in introducing nonpermeable compounds. However, the incorporation of some substances as substrates of low molecular weight and macromolecules as proteins and nucleic acids constitutes the base of some developing applications. Cells permeabilization involves two stages: induction of heterogeneities in the membrane structure and transport of molecules through these 'gates'. For great part of biotechnological applications, the process must be reversible once cell viability should be ensured. However, if the interest is some cellular component, keeping reversibility is not important [34]. Permeabilized cells are useful for analysis of complicated enzymatic processes, as protein and DNA synthesis [35]. The process of extraction and purification of enzymes is highly destructive regarding the integrity of enzymatic complexes and the intermolecular interaction. Permeabilization allows the study in situ of enzymatic activity, which better reproduces the conditions in living cells. Enzymes are kept in physiological concentration inside cells, making possible the reconstitution of metabolic paths [36]. Cordeiro and Freire (1995) [36] developed a technique of permeabilization of Saccharomyces cerevisiae that was very efficient for assays in situ of total enzymatic activity. Although in this type of assay the activities are express in terms of cells number, a more suitable base is the amount of present protein. The same authors developed latter a method that allows the protein determination in permeabilized cells. This method is based on the fact of that, after permeabilization, intracellular proteins become accessible the reagent used in the analysis, with the advantage of involving neither cells rupture nor proteins solubilization. The best permeabilization procedure depends on the organism and the composition of the membrane and cell wall. The type of reaction to be investigated is also important [35]. The chemical method is based on chemical modification of the cell membrane for substances that modify its structure. They are cheap and easy to use but also less specific and result in extensive damages to the membrane or even loss of cells viability. Detergents are the chemical agents frequently used. They act solubilizing some components of the membrane [34]. Examples of used detergents are Tween 80, Triton X, sodiumdodecyl sulphate (SDS), cetyltrimethylamonium bromide (CTAB). Among solvent organic, toluene is the most frequently used [35].
248
Biocatalytic Membrane Reactor With Continuous Removal Of Organic Acids By Electrodialysis - Borges
5.1
Zymomonas mobilis permeabilization
The cell wall, present in Gram-negative bacteria, protects the cell and maintains enzymes inside the periplasmic space. A trimeric protein in the most external layer of the cell wall is responsible for its porosity and for the diffusion of low molecular weight molecules, as sugars, for periplasm. Differences between Gram-positive and Gram-negative bacteria must be taken in account during permeabilization. For the latter, successful permeabilization requires the agent to penetrate the cell wall and reach the cytoplasmic membrane [35]. Enzyme glucose-fructose oxidoreductase present in Zymomonas mobilis promotes simultaneous oxidation of glucose to glucono-7-1actone and reduction of fructose to sorbitol. In a subsequent reaction, glucono-7-1actone is hydrolyzed to gluconic acid [28]. In intact cells, gluconic acid is converted to ethanol while sorbitol accumulates in the medium, as shown in Figure 1. When cells are permeabilized, metallic ions and cofactors can diffuse out of the cell, so that gluconic acid conversion to ethanol cannot be achieved. Because the NADP(H) cofactor, necessary for GFOR activity, remains tightly bound to it, this enzyme activity is preserved [37].
SORBITOL --
NADI~.,
~ FRUCTOSE' /
GLUCOSE
os~4arcmse ~t~ ~~ go.ludemer/ -N/kDP
k M GLUCOltOLACTONA
~ ~11ucoao lactonase GLIICONATE• H+
ATPADP ~
9t~u,noki~s~
6-PHOSPHOGLUEONATE
via Enlner-Doudoroff pathway
ETHANOL
Figure 1" Production of sorbitol and gluconic acid by Zymomonasmobillis. After permeabilization, the pathway from gluconate to ethanol, delimitedby broken line, is not functional anymore. 249
Biocatalytic Membrane Reactor With Continuous Removal Of Organic Acids By Electrodialysis - Borges
For Zymomonas mobilis permeabilization, toluene [28], CTAB [29] and freeze-thaw method [38] have been used. Figure 2 shows the effect of permeabilization on Zymomonas mobilis cells.
(A)
(B)
Figure 2: Electron transmissionphotomicrograph of Zymomonas mobilis cells before permeabilization (A) and after permeabilization (B). Besides sorbitol and gluconic acid, Zymomonas mobilis can be used to obtain many other products, such as ethanol and levan. The presence of the enzyme GFOR allows a utilization of an extensive range of substrates. From monosaccharides and disaccharides like galactose and lactose is possible to produce galactonic and lactobionic acids, respectively, and sorbitol [39]. Recombinant DNA technique can be used to produce 13-carotene, D-alanine and acetaldehyde [40]. 5.2
Immobilization of Zymomonas mobilis A few supports have been used for Zymomonas mobilis immobilization for sorbitol and gluconic acid production: calcium alginate [28] ~c-carrageenan [29] [41] and the use of a flocculent strain [42]. Ro and Kim [43] used a system where invertase was co-immobilized with cells, making possible the use of sucrose, a much cheaper substrate, instead of a solution of glucose and fructose. Another important issue to be considered is down stream treatment of the solution containing the products and some unreacted substrates. For this particular system, electrodialysis has the advantage of easy scale up and not introducing external compounds to medium. If an aqueous solution containing sorbitol and gluconic acid is submitted to an electrical potential, gluconate ions will migrate across the anion exchange membrane toward the anode compartment. Likewise, hydrogen ions will migrate through the cation exchange membrane to the cathode compartment. This allows
250
Biocatalytic Membrane Reactor With Continuous Removal Of Organic Acids By Electrodialysis - Borges
pH control without addition of a neutralizer. Sorbitol, an uncharged molecule, remains in the medium. In our work, a hollow fiber bioreactor was coupled to an electrodialysis unit, resulting in separation of gluconic acid produced by permeabilized cells of Zymomonas mobilis. 6.
EXPERIMENTAL
6.1
Microorganism and culture conditions All experiments were performed with Zymomonas mobilis CP4 (ATCC 31821), grown in a medium containing 100 g glucose/L and 5 g yeast extract/L at 30 °C. 6.2
Permeabilization Cells were permeabilized by addition of cetyltrimethylamonium bromide (CTAB) to the cells in the medium after about 20 h growth, to give a final proportion of 0.04 g CTAB/g cells. After gently stirring for 30 min, cells were separated by centrifugation at 5,000 rpm and washed twice with distilled water. 6.3
Immobilization method
6.3.1 Entrapment in alginate: For immobilization in calcium alginate, a concentrated cells suspension was mixed with a solution of sodium alginate (8 % w/v) [44] and spherical beads were produced by dropping the mixture into a 20 g/L calcium chloride solution through a needle 0.5 mm diameter. The mean diameter of obtained beads was ca. 2 mm. Beads were suspended in a 0.5 % (v/v) glutaraldehyde solution under stirring for 30 min. After washing with water, the beads were stored at 4 °C until further use. This procedure ensures that a soft gel is produced, what is important to minimize mass transfer resistance in this system. 6.3.2 Confinement in hollow fiber: Permeabilized cells suspension was confined either in the bore of microporous polycarbonate hollow fibers or in the shell side of the module. These fibers, prepared by phase inversion technique using spinning conditions reported in the Literature, were about 20 cm long, had outer and inner diameters of 0.68 and 0.33 mm, respectively and were assembled in a longitudinal module of PVC as shown in Figure 3(A). Figure 3(B) shows a photomicrograph of a fiber cross-section. The set up of the membrane bioreactor was completed by filling either the bore side of the fibers or the shell side of the module with a cell suspension. The pore size in the fiber outer surface is smaller than 0.1 0m, while in the inner surface is smaller than 1
251
Biocatalytic Membrane Reactor With Continuous Removal Of Organic Acids By Electrodialysis - Borges
~tm. This range of pore size in the fiber is small enough to avoid cells loss during experiments. Water permeability of these fibers was 912 L h - l m -2 bar. TI
__~ ......
I
~
,....... ..... ~ ]
~
Cells in the bore
~
Fibers cross section Substrates
Ceils in the shell side
...........................
(A)
(B) Figure 3 - Hollow fiber module used to confine the permeabilized cells. (A) Scheme of the hollow fiber bundle inside a module; (B) Photomicrograph of the cross section of a hollow fiber.
6.4
Assays and analytical methods
Cells concentration was determined by optical density at a wavelength of 600 nm. Protein assays were done according to the binding method of Bradford slightly modified. Glucose, fructose, sorbitol and gluconic acid concentrations were analyzed by HPLC (Waters, model 510) with refractive index detector and a Polyspher CH CA column. The column temperature was kept at 80 °C and water was used as eluent at a flow rate of 0.5 mL/min
6.5
Electrodialysis Unit
The ED stack contained three acrylic compartments, separated by two ion exchange membranes, 3,5 mm distant one of the other. Each membrane was 6
252
Biocatalytic Membrane Reactor With Continuous Removal Of Organic Acids By Electrodialysis - Borges
cm of diameter, totalizing an effective area 56 cm2. Commercially available membranes named CR 67 HMR-412 (cationic exchange membrane) and AR 204 SZRA-412 (anionic exchange membrane), from the Ionics Inc. were used. A centrifugal pump with flow control by means of a power supply was used for continuously circulate the reaction medium through the feed chamber (ca. 20 mL), at flow rate of 12-24 L/h. A 20 g/L NaC1 solution was used in the other two compartments of 800 mL each, for the nickel electrodes rinse. This way, in our study gluconate was recovered as a sodium salt, although gluconic acid can be the final product if a proper assemble of the ED stack is used. The voltage and the current across the ED unit were controlled by a D. C. power supply. The figure 4 show the removal rate of gluconic acid as a function of applied voltage. To avoid water splitting the voltage in electrodialysis unit coupled reaction experiments was fixed at 20 V, which corresponds to a current density always lower than 5 mA/cm2. System temperature was controlled at 39 °C. 0.3 At "
E
0.25
01 q)
t~
0.2
0
E 0.15 C ..,,,
0.1 10
15
20
25
30
ddp (V)
Figure 4 - Removal rate of gluconic acid as a function of applied voltage. 6.6
Kinetic studies Experiments were operated in batch by adding free cells or calcium alginate immobilized cells to the reaction medium or by circulating an equimolar fructose/glucose solution through the hollow fiber bioreactor. Temperature was constant at 39°C as well as the medium pH, kept constant at 6.2 by automatic titration using 1 M NaOH when in absence of the electrodialysis unit. In the coupled system, the setup is shown in Figure 5.
7.
RESULTS
The initial specific reaction rate was determined in experiments carried out with cells immobilized in calcium alginate or confined in hollow fibers. 253
Biocatalytic Membrane Reactor With Continuous Removal Of Organic Acids By Electrodialysis - Borges
A/IC
in tl pJ ( 7 )
(2)
i (1)
(4)
(3) (6)
(5) Figure 5 - Experimental set up for the experiments with electrodialysis. (1) Thermostatic bath (2) Hollow fiber bioreactor (3) Pump (4) pH controller (5) Magnetic stirrer (6) Substrates tank (7) Electrodialysis stack.
Reactions using free cells were also carried out to analyze the mass transfer effect originated by the immobilized system. Table 1 shows initial specific reaction rate of gluconate production and protein mass used in each experiment. As can be seen, the system utilizing cells immobilized in calcium alginate presented the lowest initial specific reaction rate. In fact, in most gel entrapment systems, diffusion through the gel can be the rate-limiting step reducing the reaction rate. Hollow fiber confined cells, on the other hand, showed an initial specific reaction rates comparable to those using free cells. In this case, the mass transfer resistance is much lower than in the gel, because diffusion occurs in the liquid phase through the membrane pores. Table 1 - Performance of different cell immobilization methods without coupling ED unit. Immobilization method
Initial specific reaction rate (gsorbitol/gprotein.h) 32.8
Protein mass (mg) 64.5
Calcium alginate
3.8
145.7
Bore of hollow fibers
27.9
8.6
Shell side of hollow fibers module
16.5
17.4
Free cells
Note: A reaction volume of 100 ml was used in all experiment
254
Biocatalytic Membrane Reactor With Continuous Removal Of Organic Acids By Electrodialysis - Borges
The specific initial rate for the cells confined in the shell side of the module was about 40% lower than that for cells in the bore of the fibers, which can be attributed to a less efficient pH control due to cell deposition occur in the shell of the module. It should be noted that, although different protein mass was used in each experiment, these results qualitatively indicate that confined cells can lead to reaction rates very close to that obtained with free cells. 7.1
Coupled System
7.1.1 Cells confined in the shell side of the hollow fiber module Reaction medium was continuously pumped through the electrodialysis unit at a flow rate of 24 L/h (Reynolds number of 1230). Gluconic acid was continuously removed, keeping its concentration in the medium lower than 1 g/L. Despite gluconic acid removal, the pH of the reaction medium was about 5.5, a little far from that ideal for GFOR activity (6.2). Reaction evolution was accompanied by sorbitol production. Initially, a protein mass of 36 mg for 100 ml of reaction volume was confined in the shell side of the module. For this experiment, the specific reaction rate remained constant at about 2.6 gsorbitol/gprotein.hduring more than 25 hours. When more cells were confined in the module, corresponding to a protein mass of 73.5 mg for 100 ml of reaction volume it was observed a decrease of the specific reaction rate, as shown in Figure 6. This fact can be attributed to local reduction of pH, which becomes more critical at high concentration of cell suspension. Recalling results presented in Table 1 (rate of 16.5 gsorbitol/gprotein.h) for the experiment at same configuration but without electrodialysis, also reinforces this conclusion. It should be noted that, although different protein mass have been applied in each experiment, the observed reduction in the effective reaction rates only can be justified by unfavorable conditions for the enzymatic reaction. In this way, it becomes evident the necessity of module design optimization to achieve ideal conditions for product removal and pH control. 7.1.2 Cells confined in the bore of the hollow fiber module In order to improve product removal, cells were confined in the bore of the fibers, where the stagnant layer is much thinner and the local concentration is better controlled. At this time, the flow was kept at 12 L/h (Reynolds number of 185) to minimize the water flux through ion exchange membranes. A lower cell mass was used, corresponding to a protein concentration of 16 rag/L, as consequence of the reduced volume inside fibers. Figure 7 shows the result of the experiments. A specific reaction rate of 17.5 gsorbitol/gprotein.hwas observed in this experiment, which is 4 times higher than that obtained in the system using
255
Biocatalytic Membrane Reactor With Continuous Removal Of Organic Acids By Electrodialysis - Borges
80
30
70 25
tO
(n o
60 A
t..4) ¢,1 C
0 o o o 4) e~ 0 ..Q L_ 0
._.R m ,4-* e
o-
-20
50
~: m (,,) O
40
-15
~i'D
-lO
g N"
o 30
O j e ¢ 1///'I/j/
0 1/)
.~
20
-5
10
,
t
0
5
10
15
20
o
25
3o
T i m e (h)
Figure 6 - Cells confined in the shell side of the hollow fiber module. Effect of the protein concentration on the specific reaction rate: ( 0 ) 36.0 mg of protein, ( e ) 73.5 mg protein. Sorbitol specific concentration ( ~ ) and sorbitol concentration ( .... ) evolution. Initial concentration: Glucose = 100 g/L, Fructose = 100 g/L.
700
16 /e
0 .,.=
E
A
t41
._¢ 41
r. 0
,_ Q.
o
600
-
500
-
400
-
0 /
-14 -12
O -I
o" O
-10
n 0 ¢.)
-
e~ ._ 0 L 0 t~
L. 0 w
300
8 :3 .ll
-
200
i ./
~,.,o
-6
1/
0
/ /
4
ca
100 0 ~ 0
. . . . . . . . . . . . . 5 10 15 20 25 Time
0 30
35
(h)
Figure 7 - Cells confined in the bore of the hollow fiber module. Effect of the protein concentration on the reaction rate: (e) One module, initial protein mass of 2,2 mg, (n) Two modules in series, initial protein mass of 4,4 mg. Sorbitol specific concentration ( ~ ) and sorbitol concentration ( .... ) evolution. Initial concentration: Glucose = 100 g/L, Fructose = 100 g/L.
256
Biocatalytic Membrane Reactor With Continuous Removal Of Organic Acids By Electrodialysis - Borges
cells confined in the shell side of the module. When two identical modules in series were used, duplicating the total protein mass, the specific reaction rate was not significantly different from that using one single module. It demonstrate that an increase in reaction rate proportionally to the protein concentration used can be achieved by distributing the total cell mass among distinct modules, which avoids high local cells concentration. Another remarkable conclusion from these experiments is related to the stability of this system. The same hollow fiber module was used for more than 64 hours, with no noticeable decrease of the enzymatic activity. A very different behavior was observed in a system without gluconic acid removal by electrodialysis. In this case, it was verified a decrease of the enzyme activity in the initial stages of the reaction. To discard the possibility that the consume of substrates have caused the observed decrease in the reaction rate, after about 100 h of operation reaction medium was replaced by a fresh substrate solution. As shown in Figure 8, the specific reaction rate for the new substrate solution was almost the same as that observed in the final stages of the previous reaction. This indicates that enzyme inactivation, rather than a substrate starving system, is responsible by the decrease of the enzyme activity. 127 _.
9
o ,,a I
m o
mini
o m
~ n,- 6
6
¢D
o .'~'o -1 A
f,Q
3
I"
l" 0
'
I 20
'
I 40
' Time
I 60
'
I 80
'
I 100
(h)
Figure 8 - Experiments without electrodialysis coupling. Specific reaction rate for the first experiment ( ~ ) and after substrate solution replacement (---). Sorbitol concentration (+). Initial concentration: Glucose = 100 g/L, Fructose = 100 g/L.
257
Biocatalytic Membrane Reactor With Continuous Removal Of Organic Acids By Electrodialysis - Borges
Figure 9 is a comparison between specific reaction rates for this experiment and for that using electrodialysis coupling. The initial rate in the system with or without electrodialysis unit should be the same. However in the experimental coupled system it was not possible to observe the transient period. Nevertheless, a sharp increase in the reaction rate during the early stages is expected. It is clear that enzyme stability improved when gluconic acid was removed by electrodialysis. The presence of gluconic acid causes an inhibition of the lactone hydrolysis and its accumulation in the reaction medium. Interaction with the lactone is responsible for GFOR inactivation [45]. In this way, the removal of gluconate ion by electrodialysis has a favorable effect once it prevents the lactone accumulation. 16
12
~.__. .2 o
4
0
0
10
20
30 Time
40
50
60
70
(h) Figure 9: Specific reaction rate for the system with (---) and without ( ~ ) electrodialysis. 8.
CONCLUSIONS
The electrodialysis unit coupled to the reactor allowed an efficient removal of gluconic acid. Its concentration in the reaction medium was kept lower than l g/L. Furthermore, it was observed an improvement in the stability of the enzyme. It was not found any reduction in the reaction rate, even after 60 hours of reaction. On the other hand, when NaOH was applied to neutralize the gluconic acid produced, it was verified a reduction of 80 % in the reaction rate in the same period.
258
Biocatalytic Membrane Reactor With Continuous Removal Of Organic Acids By Electrodialysis - Borges
The use of permeabilized cells in the bore of hollow fibers shown the best performance in terms of specific reaction rate. The mass transfer resistance in this case was much reduced when compared to that of cells immobilized in alginate beads. However, if cells are confined in the shell side of hollow fiber module it was find a lower enzyme activity, which was attributed to local pH reduction due to cell deposition. This fact leads to the conclusion that the balance between cell concentration conf'med by hollow fibers and local pH is a very important factor for the design of this kind of reactor. The successful obtained with the studied case might be extended to similar reaction systems. The coupling of an electrodialysis unit to the bioreactor showed to be an efficient method to avoid acid inhibition effects, as well as to control the pH of the reaction medium. On the other hand, the confinement of cells by microporous hollow fibers facilitates the scale-up of the bioreactor and, in comparison to conventional immobilization methods, it allows to operate at low mass transfer resistance leading to elevated reaction rates. REFERENCES [1]
[21
[3]
[4] [5] [6]
[7] [8] [9] [10] [11] [12] [13]
M. Seki, S. Furusaki, Shigematsu, Kunihiko, Cell Growth and Reaction Characteristics of Immobilized Zymomonas mobilis, Ann. N. Y. Acad. Sci. 613 (1990) 290-302. E. Drioli and L. Giorno, Biocatalytic Membrane Reactors, Taylor & Francis Ltd, Padstow, UK, 1999. H. A. Fishman, D. R. Greenwald, R. N. Zare, Biosensors in Chemical Separations, Annu. Rev. Biophys. Biomol. Struct. 27 (1998) 165-198. A. M. Vaidya, P. J. Halling and P/J. Bell, Novel Bioreactor for Aqueous-Organic Two-Phase Biocatalysis: A Packed Bed Hollow-Fiber Membrane Bioreactor, Biotech. Tech. 7 (1993) 441-446. S. Doig, A. T. Boam, D. I. Leak, D. C. Stuckey, A Membrane Bioreactor for Biotransformation of Hydrophobic Biomolecules, Biotech. Bioeng. 58 (1998) 587594. D. M. F. Prazeres and J. M. S. Cabral, Enzymatic Membranes Bioreactors and their Applications, Enzyme Microb. Technol. 16 (1994) 738-750. M. Gramer, D. M. Poeschi, M. J. Conroy, B. E. Hammer, Effect of harvesting Protocol on Performance of a Holow Fiber Bioreactor, Biotch. And Bioeng. 65 (1999) 335-340. H. Nakano, T. Shiinbata, R. Okumura, S. Sekiguchi, M. Fujishiro, T. Yamane, Efficient Coupled Transcription from PCR Template by a Hollow-Fiber Membrane Bioreaetor, Biot. Appl. Bioeng. 64 (1999) 194-199. A. W. Brunch, The Uses and Future Potential of Microbial Hollow-Fibre Bioreactors, J. Microbiol. Methods 8 (1988) 103-119. G.B Borglum, and J.J. Marshall, Appl. Biochem. Biotechnol. 9 (1984) 117-130. J. Vaija, Y.Y. Linko, P. Linko, Appl. Biochem. Biotechnol. 7 (1982) 51-54. H. Eikmeier and H. J. Rehm, Appl. Microbiol. Biotechnol. 20 (1984) 365-370. M. Cheryan and M. A. Mehaia, Continuous fermentation with membrane bioreactors, 4 th Int. Congress Eng. Food (1985) Alberta, Canada.
259
Biocatalytic Membrane Reactor With Continuous Removal Of Organic Acids By Electrodialysis - Borges
[14] [15] [16] [17] [18]
[19]
[20] [21] [22] [23] [24] [25] [26] [27]
[28]
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[32] [33]
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M. A. Mehaia and M. Cheryan, Production of lactic acid from sweet whey permeate concentrates, Process Biochemistry 22 (1987) 185-188. H. Ohara, K. Hiyama, T. Yoshida, Kinetics of growth and lactic acid production in continuous and batch culture, Appl. Microbiol. Biotechnol. 37 (1992) 544-548 H. Ohara, K. Hiyama, T. Yoshida, Lactic acid production by a filter-type reactor, J. Ferment. Bioeng. 76 (1993) 73-75. Y. Shimizu, K. Shimodera, A. Watanabe, Cross-flow filtration of bacterial cells, J. Ferment. Bioeng., 76 (1993) 493-500. V. Visacky, H. Shtrathmann, E. Drioli, A. Narebska, F. P. Cuperus, S. Schlosser, Spacer for hybrid and multistage membrane process, Slovak patent application PV 0691-1996, op. cit. in Drioli and Giomo, 1999. P. Boyaval, P. and C. Corre, Continuous lactic acid fermentation with concentrated product recovery by ultrafiltration and electrodialysis, Biotechnol. Lett. 9 (1997) 207212. S. G. M. Van Nispen and R. Jonker (1991) U.S. Patent 5,002,881 A. Narebska, Second Annual Report of Copernicus Project, ERB-CIPA CT92 3018, Op. cit. in Drioli and Giomo, 1999. P. Boyaval and C. Corre, Production ofpropionic acid, Boyaval P, Corre C, Lait. 75 (1995) 453-461. L. R. Schlicher and M. Cheryan, Reverse osmosis of lactic acid fermentation broths, J. Chem. Technol. Biotechnol. 49 (1990) 129. L. Giorno, P. Spicka, E. Drioli, E., Downstream processing of lactic acid by membrane based solvent extraction, Sep. Sci. Technol. 31 (1996) 2159-2169. Juang and Huang, Kinetic studies on lactic acid extraction with amine using a microporous membrane-based stirred cell, JMS 129 (1997) 185-196. Y, P Tong, M. Hirata, H. Takanashi, T. Hano, Back extraction of lactic acid with microporous hollow fiber membrane, JMS 157 (1999) 189-198. M. Zachariou and R. K. Scopes, Glucose-Fructose Oxidoreductase, a New Enzyme Isolated from Zymomonas mobilis that is Responsible for Sorbitol Production, J. Bacteriol. 167 (1986) 863-869. H. Chun and P. L. Rogers, The Simultaneous Production of Sorbitol from Fructose and Gluconic Acid from Glucose Using an Oxidoreductase of Zymomonas mobil& Appl. Microbiol. Biotechnol. 29 (1988) 19-24. B. Rehr, C. Wilhelm, and H. Sahm, Production of Sorbitol and Gluconic acid by permeabilized cells of Zymomonas mobilis, Appl. Microbiol. Biotecnhol. 35 (1991) 144-148. S. L. Paterson, A. G. Fane, C. J. D. Fell et al., Sorbitol and Gluconate Production in a Hollow Fibre Membrane Reactor by Immobilized Zymomonas mobil&, Biocatalysis, 1 (1988) 217-229. M. Matey, The Production of Organic Acids, Critical Rev. Biotechnol. 12 (1992) 87132. K. Q. Wilberg, T. L.M. Alves and R. Nobrega, Enzymatic Catalysis by Permeabilized Cells, Brazilian J. Chem. Eng. 14 (1997) 347-352. H. C. Ferraz, C. P. Boreges, T. L. M. Alves, Coupling of an Electrodialysis Unit to a Hollow Fiber Bioreactor for Separation of Gluconic Acid from Sorbitol Produced by Zymomonas mobil& Permeabilized Cells, JMS 191 (2001) 43-51. I. Hapala, Breaking the Barrier: Methods for Reversible Permeabilization of Cellular Membranes, Critical Rev. Biotechnol. 17 (1997) 105-122. H. Felix, Permeabilized Cell, Anal. Biochem. 120 (1982) 211-234.
260
Biocatalytic Membrane Reactor With Continuous Removal Of Organic Acids By Electrodialysis - Borges
[36] [37]
[38] [39]
[40] [41]
[42] [43]
[44] [45]
C. Cordeiro and A. P. Freire, Digitonin Permeabilization of Saccharomyces cerevisiae Cells for in Situ Enzyme Assay, Anal. Biochem. 229 (1995) 145-148. J-P Park and H-S Kim, A New Biosensor for Specific Determination of Glucose or Fructose Using an Oxidoreductase of Zymomonas mobilis, Biotechnol. Bioeng. 36 (1990) 744-749. S. Bringer-Meyer and H. Sahm, Process for Obtaining Sorbitol and Gluconic Acid by Fermentation and Cell Material Suitable for this Purpose (1991) US Patent 5,017,485. M. Satory, Continuous Enzymatic production of Lactobionic Acid Using GlucoseFructose Oxidorreductase in an Ultrafiltration Membrane Reactor, Biotechnol. Letters 19 (1997) 1205-1208. H. W. Doelle, L. Kirk, Crittenden et al, Zymomonas mobilis- Science and Industrial Application, Critical. Rev. Biotechnol. 13 (1993) 57-98 K. H. Jang, C. J. Park and U. H. Chun, Improvement of Oxidoreductase Stability of Cethyltrimethylammoniumbromde Permeabilized Cells of Zymomonas mobilis Throug Glutaraldheyde Crosslinking, Biotechnol. Letters 14 (1992) 311-316. E. Wisberg, Evaluation of the Flocculent Strain Zymomonas mobilis Z 1-81 for the Production of Sorbitol and Gluconic Acid, J. Basic Microbiol. 6 (1997) 445-449. H. S. Ro and H. S. Kim, Continuous Production of Gluconic Acid and Sorbitol from Sucrose Using Invertase and an Oxidoreductase of Zymomonas mobilis, Enzyme Microbiol. Biotechnol. 13 (1991) 920-924. J. E. Bailey and D. F. Ollis, Biochem. Eng. Fundamentals, McGraw-Hill, New York, NY, 1986. M. F~linger, D. Haltrich, D. Kulbe et al., A Multistep Process is Responsible for Product-Induced Inactivation of Glucose-Fructose Oxidoreductase from Zymomonas mobilis, Eur. J. Biochem. 251 (1998) 955-963.
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New Insights into Membrane Science and Technology: Polymeric and Biofunctional Membranes D. Bhattacharyya and D.A. Butterfield (Editors) © 2003 Elsevier Science B.V. All rights reserved.
Chapter 13
Use of micro-porous affinity membranes for protein purification: a case study F. Cattoli and G. C. Sarti*
DICMA - Dipartimento di Ingegneria Chimica, Mineraria e delle Tecnologie Ambientali, Alma Mater Studiorum- Universit~ di Bologna, viale Risorgimento 2 40136 Bologna, Italy *Corresponding author, Tel.: +39.0512093142, Fax: +39.051581200, e-mail: [email protected] ABSTRACT The recovery of MBP-intein-CBD using amylose affinity membranes has been studied. MBP-intein-CBD is a model fusion protein containing three different protein domains: MBP as the affinity target, CBD as the protein of interest and the intein domain inserted between MBP and CBD in view of its self-cleavage activity. Such a fusion protein allows for the recovery of the target product in a single chromatographic step. The affinity separation process is conducted exploiting the specific interaction of the MBP group towards amylose. The use of micro-porous membranes as stationary phase is considered in order to improve the fluiddynamic of the system, i.e. flow rate, pressure drop and mass transport in the liquid phase. A large scale membrane holder has been used, suitable for the separation of large amounts of the desired product. Good results in terms of purity and concentration of the product were achieved. Kinetic and equilibrium data have been obtained; the experimental equilibrium curve was well represented by the Langmuir isotherm; the kinetic constants were determined based on a simplified kinetic model. 1.
INTRODUCTION
Downstream processes play a major role for the recovery of biomolecules from complex mixtures; often they can account for more than the 50 % of the total costs of the entire production process. Recovery of biomolecules is a challenge due to the restricted conditions (temperature, pH, ionic strength, solvents) suitable to perform the separation process. One of the requirements is a reduced sequence of separating steps together with an excellent yield in terms of concentration and purity of the recovered 263
Use Of Micro-Porous Affinity Membranes For Protein Purification: A Case Study - Sarti
product. Several chromatographic techniques have been developed and are currently adopted such as ion exchange chromatography, immobilized metal affinity chromatography, hydrophobic interactions, gel filtration chromatography and others. Among all these chromatographic techniques, affinity chromatography is proposed as one of the finest, due to the separation mechanism exploited. The desired product is selectively removed from a complex mixture thanks to a highly specific interaction occurring between the active binding site of the biomolecule and a complementary substance, specifically recognized. Usually the ligand is bound to an insoluble polymeric surface. Stationary phases for affinity chromatography processes are often obtained and produced in form of resins or porous beads, packed in column configurations. As already discussed by many authors, the fluid-dynamics of these systems offers several limitations such as diffusion mass transfer control and high pressure drops, with consequent low flow rates attainable. All these factors combined together lead to the long time typically required to complete a single separation step. The use of micro-porous supports, such as microfiltration membranes, has been proposed in order to overcome the main fluiddynamic limitations offered by resins. The open porous structure of the matrix allows the liquid solution to flow over the binding surface: the biomolecules are, in this way, conveyed towards the binding sites where the immobilization reaction occurs. There is, thus, a strong reduction of the processing times thanks to the reduced contribution of diffusion on the overall mass transfer rate. Considering only geometric factors, the diffusion paths in micro-porous membranes are of the order of the pore radius while in porous spherical beads the diffusion paths depend on the dimensions of beads and thus are at least one order of magnitude higher. Micro-porous membranes offer, on the other hand, an unfavorable ratio of the total external surface per unit volume of stationary phase. Nevertheless, very high concentration of the ligand in the beads are not desirable in order to reduce problems connected with multiple-site binding. Thus, when the diameter of membrane pores is small enough, the surface area available for ligand immobilization on micro-porous membranes is approximately equivalent to that offered by spherical beads. Hence, compared with the high variety of affinity column materials already available, the advantages achievable by affinity membranes are due to kinetic effects, when all other equilibrium properties are equivalent. In fact, by speeding up the diffusion process, the controlling step becomes the rate of complex formation. One of the great advantages of affinity membranes compared to affinity resins is the increase in the production rates due to extended bed configuration and reduction of diffusion resistances [ 1-3]. In this work we propose the use of affinity micro-porous membranes for the recovery of a class of fusion proteins, characterized by the presence of the 264
Use Of Micro-Porous Affinity Membranes For Protein Purificatiot: A Case Study - Sarti
Maltose Binding Protein (MBP) domain. The MBP is a periplasmic protein involved in the transport of maltose throughout the cell membrane. This MBP domain is selective in binding maltose, cyclodextrins and other long chain polysaccharides like amylose [4-5]. Fusion based affinity systems can be seen as simple means for the recovery of a desired product fused to a peptide, e.g. the MBP group, which is selectively bound to a complementary ligand. MBP fusion systems are commercialized by New England Biolabs (Beverly, MA), and two different types are available: the pMAL Protein Fusion System [6] and the IMPACT-CN system [9]. Conventionally, the DNA sequence coding for the protein of interest is fused to the DNA sequence of a binding protein to form a single open reading frame. The plasmid vector obtained expresses a two-domain fusion protein that can be easily purified exploiting the specific interaction between the binding domain and its immobilized ligand. In the pMAL system the two different protein domains are joined through a recognition sequence of a protease, in this case Factor Xa [6-8]. Once purified through a specific affinity system, the two protein domains are cleaved off thanks to the action of Factor Xa. The product solution contains, at this point, the binding domain cleaved from the target product and the amount of the protease added; the MBP domain may also be in solution, if cleavage is performed after elution. Subsequent purification stages allow the removal of the protease and of the MBP groups, if any, to obtain the isolated purified product. Some disadvantages in the use of this system can be due to non specific cleavage, costs connected with the use of more protease as the scale of the system increases and finally the possible need of an additional purification passage. The IMPACT-CN system offers an advantage compared to the pMAL system due to the introduction of self-cleaving protein linker. In particular, for affinity separations a binding domain is fused to a self splicing protein element, as intein, which is in turn fused to the target protein. Once the fusion has been separated from the complex protein mixture, the self-cleavage reaction allows to obtain the target protein; if the self-cleavage reaction is carried out on the fusion protein still loaded on the column, no further purification step is needed. Intein has been modified so that it undergoes a self-cleavage reaction at its Nterminus at low temperature, in presence of either dithio-threitol or [3mercaptoethanol or cysteine. Other authors have obtained a different smaller intein able to undergo a C-terminus splicing reaction in precise operative conditions [ 10]. Our experiments have been conducted by using MBP-intein-CBD (i.e. Chitin Binding Domain) fusion protein obtained by genetically modified E. coli strains (New England Biolabs) [9]. In particular, the selectivity of amylose affinity membranes has been tested performing the purification process directly from cell lysate; no specific attention was paid to the splicing reaction. Pure 265
Use Of Micro-Porous Affinity Membranes For Protein Purification: A Case Study - Sarti
protein solutions have then been used in adsorption experiments, to obtain equilibrium and kinetic data.
2.
THEORY
2.1
Adsorption Isotherm The model used to describe the adsorption isotherm in batch systems considers a "one to one" interaction between the protein and the ligand immobilized onto the matrix as in a mono-layer adsorption mechanism: P
(protein)
+
L
( ligand )
(1)
+-~ PL
(complex)
According to a very common procedure, the mass balance equation for the solute adsorbed on the solid phase can be written, at any time t, by considering the kinetics of forward and backward reactions. The latter is usually considered to be first order in the surface protein concentration; for the former a linear dependence on the protein concentrations both in solution and on the surface is considered. The kinetic equation thus is:
(2)
dq = k, (q, n - q)c - kzq
dt
where q is the protein surface concentration at time t (mass of protein per unit membrane area), qm is the maximum binding capacity of the affinity support, c is the protein concentration in the liquid solution, k/is the rate constant of the forward reaction and k2 is the rate constant of the backward reaction. At equilibrium one immediately obtains the well known Langmuir isotherm: q = qm-------f----C K d +c*
(3)
where:
Kd
k2 = k---~-
(4)
The symbols q* and c* represent the equilibrium values of the protein concentration on the surface and in the liquid solution, respectively, and Ka is the equilibrium constant of the binding reaction.
266
Use Of Micro-Porous Affinity Membranes For Protein Purification: A Case Study = Sarti
2.2
Short Times Analysis. Considering only the initial instants of each adsorption experiment, some further simplifying assumptions can be introduced. At the early stages of adsorption, the general kinetic expression (2) can be simplified retaining only the forward reaction and neglecting the backward reaction term, since the latter becomes effective only after an appreciable amount of protein has been loaded. To strengthen the validity of this assumption, we notice that the backward reaction is indeed rather slow during the adsorption step and needs the use of an appropriate elution solution to become appreciable. At the same time, at the early stages of adsorption, the protein concentrations in solution and on the membrane surface can be taken as their respective initial values, so that c--->Coand (qm-q)--->qm; one thus obtains:
(5)
dq -__k~q,Co dt
This equation coupled with the mass balance for a batch system (6)
V --dc = - A dq dt dt
and integrated over time leads to the final expression: (7)
c = c o -aklqmCot
In the above equations, Co is the starting protein concentration in solution, V is the solution volume and A is the area of the membrane sheet. The quantity a, defined as A/V, is a working parameter. Equation (7) can be reasonably applied to interpolate only the initial adsorption data; therefore, in nmning the experiments the time interval between two subsequent data points must be sufficiently small, in order to have proper comparison with the simplified model above [11]. Similarly, the early stages of elution can be described by equation (2) in a simplified way by considering only the backward reaction term. By taking the initial value of q as q*, equation (2) thus becomes: (8)
dq__ k2,q , dt
Coupling equation (8) with equation (6) and integrating over time the following expression is obtained: 267
Use Of Micro-Porous AtYmity Membranes For Protein Purification: A Case Study - Sarti
(9)
c : ak2' q * t
where q* represents the amount of protein immobilized on the membrane surface at the end of the adsorption step and k2' is the binding constant of the desorption reaction, under elution conditions. Of course this value differs from the k2 value holding true in the adsorption conditions. 3.
MATERIALS AND METHODS
3.1
Affinity Membrane Preparation Amylose affinity membranes are obtained through the chemical modification of a native cellulose matrix. Cellulose is mechanically and chemically stable, is biologically inert and, in view of its hydrophilicity, shows low not specific interactions. In this work Whatman 541 membranes have been used. The main features of the membranes adopted are reported in Table 1. Table 1. Membrane features. Membrane Dp Thickness Whatman 541
25 pm
150 pm
Porosity
Filtration speed"
Physical features
0.55
100 ml in 12 s
Hardened, ashless
* using a filter of 15 cm diameter
Amylose has been chosen as ligand due to the feasibility of the coupling reaction with the native matrix. The modification protocol is schematized as a sequence of three subsequent steps. The first step is the activation of the native support through the immobilization of a spacer arm, endowed with two terminal epoxy groups. In particular, 1,4 butanediol diglycidil ether is reacted with t h e OH groups present on the membrane surface. The reaction is conducted in a basic medium (pH>13) in presence of NaBH4 as catalyst. The second step is the immobilization of the ligand to the matrix via the spacer arm. The reaction is conducted in a mixture containing a high concentration of amylose, pH>13, in presence of NaBH4 at 37 °C. Finally, the unreacted epoxy groups after treatment with amylose solution, are blocked by soaking the functionalized membranes in a solution of mono ethanolamine, pH-9.5, at 25 °C. Amylose affinity membranes can be stored in water or in standard buffer solution at 4 °C until use. No degradation of the support has been observed over a period of 6 months [12-14].
3.2
Protein Concentration Measurement MBP-intein-CBD (97 kDa) concentration was measured performing the Bradford assay. Bradford reagent from Sigma was used. A calibration curve was
268
Use Of Micro-Porous Affinity Membranes For Protein Purification: A Case
S t u d y - Sarti
prepared by using BSA (Sigma, fraction V, 96-99%) as standard protein. The absorbance of the dye-protein complex was measured at 595 nm [ 15-16]. 3.3
Purification of MBP-Intein-CBD from Cell Lysate The fusion protein MBP-intein-CBD was obtained from genetically modified E. Coli strains ("IMPACTTM-CN '' system, from New England Biolabs Inc. ) [9, 17]. A p-TYB vector allows the fusion of the self-cleavable intcin tag to the C-terminal of the MBP group from one side and to the chitin binding domain to the other side. This procedure can be generalized by substituting the CBD domain with a different target protein. Isopropyl-[3-D-thiogalactopyranosidc (IPTG) is recognized as inductor for the expression of the fusion protein. E. Coli was grown following the conditions indicated by the manufacturer. IPTG was added during the last growing stage to stimulate the production of the recombinant protein inside the bacteria. The fusion protein is expressed as an intracellular protein. The cell culture, therefore, is harvested, the cells resuspended in buffer solution then sonicated a T = 4 °C. The cell extract is recovered, centrifuged for the removal of cell walls and debris. The sumatant, containing the protein of interest, is collected and stored at-20 °C until use. MBP-intcin-CBD is purified by a separation process conducted using amylosc affinity membranes. Flat sheet affinity membranes were located inside a membrane holder, realized in order to offer a high binding area. The total membrane area can be varied depending on the binding capacity needed to perform a specific separation process [ 18]. In particular, the membrane holder is realized with a series of discs arranged in a column configuration. On each disc a number of membranes can be located. In this work, a total membrane area of 1000 cm 2 has been used, corresponding to stacks of 3 membranes for 17 stages. The single discs arc designed in order to offer a uniform flow distribution over the cross section all along the membrane stack (Figure 1). The purification procedure is similar to the one adopted to perform the purification of other MBP-fusion proteins [19-20]. A volume of cell lysatc is diluted with buffer solution (0.5 M NaC1, 10.3 M EDTA, 0.02 M Tris HC1, pH=7.4). After a conditioning step performed on the stationary phase, the protein mixture to be purified is loaded into the circuit. The feed solution is rccirculatcd maintaining a flow rate of 12 ml/min for 80 minutes. In this condition, the specific interaction occurring between the MBP domain and the ligand present on the membrane surface allowed for the immobilization of the desired product on the membrane surface. When the adsorption step is accomplished, the stationary phase is washed with an a-specific buffer solution (0.2 M NaC1, 10.3 M EDTA, 0.02 M Tris HC1, pH=7.4) in order to remove all the fragments trapped on the membrane surface and not specifically bound. Finally, the adsorbed material is clutcd applying a buffer solution containing maltose as competing substratc for the MBP active binding site and a high 269
Use Of Micro-Porous Aft'mity Membranes For Protein Purification: A Case Study - Sarti
concentration of NaC1 (0.1 M Maltose, 0.3 M NaC1, 103 M EDTA, 0.02 M Tris HC1, pH=7.4) [21]. The product is recovered in several subsequent fractions. The protein fractions collected are then analyzed through an electrophoresis technique, using the CRITERION system from B io-Rad. Protein bands are stained using the Blue Coomassie method.
Figure 1. Large scale membrane holder, suitable for flat sheet membranes. A series of discs are arranged in a column configuration. On each discs a stack of membrane is located.
The solutions recovered during the elution stage contain an excess of unbound maltose and NaC1, besides maltose bound to the active site of MBP. A dialysis step is then performed in order to remove the excess of NaC1 as well as of maltose either free in solution or bound to the MBP domain. The conditions adopted are the same as in a previous work [22]. Several dialysis stages have been performed, lasting 24 hours each for a total time of 5 days. The system was kept at 4 °C, mechanically stirred. Pure protein solutions are thus obtained, with all the MBP active sites free of maltose.
3.4
Adsorption of Pure MBP-Intein-CBD in Batch System Pure MBP-intein-CBD solutions have been used in adsorption experiments conducted in batch system using amylose affinity membranes. In a volume of pure protein solution, variable between 10 and 14 ml, a number of amylose membranes are immersed (dm=4,7 cm). The system is kept at 4 °C and is mechanically stirred. The adsorption reaction is followed until an equilibrium condition is reached. The concentration of the protein in the liquid solution is monitored at regular time intervals, via Bradford assay and the adsorption curves are thus obtained. 270
Use Of Micro-Porous AtTmity Membranes For Protein Purification: A Case Study - Sarti
A second set of experiments has been performed loading pure MBP-intein-CBD solution, in a continuos flow mode, on a stack of 5 membranes located inside a stainless steel Millipore® membrane holder (din=4.7 cm, cross section diameter 3.8 cm). The permeate is recirculated to the stationary phase, at a flow rate of 4 ml/min, until a steady state condition is reached. The concentration in the liquid solution is monitored at regular time intervals in order to measure the kinetic of the adsorption reaction. Once a steady state condition is reached, the membranes are removed from the holder, and soaked in the recovered protein solution until a further equilibrium condition is reached. The two subsequent adsorption stages are performed at 4 °C. 3.5
Elution of Immobilized MBP-Intein-CBD in Batch System First of all, adsorbed membranes are washed with an a-specific buffer and then elution is performed by applying 20 ml of maltose buffer solution. Both stages are carried out in batch system, at 4 °C under mechanical agitation. The concentration of the released protein is monitored until a constant value is reached. 4.
RESULTS
4.1
Purification of MBP-Intein-CBD from Cell Lysate The complete saturation is achieved after 6 to 10 passages of the feed solution on the stationary phase corresponding to an adsorption time of about 80 minutes. The washing step is performed at a flow rate variable between 25 and 35 ml/min. The liquid solution is forced throughout the stationary phase without any recirculation. The efficiency of the washing step in the removal of the nonspecifically adsorbed materials is crucial in order to get a pure product at the end of the separation process. UV analysis performed on the washed-out fractions, collected at regular time intervals, shows a rapid decrease of the amount of materials dragged by the washing solution flowing out from the stationary phase. The washing step is completed in a time between 30 and 45 minutes. Electrophoresis analysis confirmed that in the last fractions collected during washing, no traces of impurities are present [20]. The elution of the specifically adsorbed material is obtained applying maltose buffer. Experimental data confirm the effectiveness of the combined action of maltose, as competing substrate, and of the ionic strength of the solution in the removal of the desired product. The recovered fractions are characterized by high purity (Fig. 2) and high protein concentrations, as showed by the elution profiles reported in Fig. 3. The two curves in Fig.3 refer to experiments performed maintaining different flow rates during the elution step. At higher flow rate, the peak of the elution profile is shifted towards higher elution volumes and the maximum concentration value is lower. This behavior can be seen as an indication of the 271
Use of Micro-Porous Affinity Membranes For Protein Purification: A Case Study - Sarti
F
E1
M
E2
E3
MBP-intein-CBD M W 97 k D a
! . . . .
Figure 2. Electrophoresis analysis of protein samples obtained from the purification process of MBP-intein-CBD using amylose affinity membranes. (F = feed, Ei = i-th eluted fraction, M -- Molecular weight Markers). 0.7 flowrate 48 ml/min ~ Exp. 1
0.6
"_ flowrate 6 ml/min [] flowrate 6ml/min ~
0.5
~
8
Exp. 2 Exp.
0.3 0.2
0.1
0
50
100
150
volume (ml)
Figure 3. Elution profiles obtained in the separation process of MBP-intein-CBD, at two different flow rates: flowrate 6 ml/min (I--!), flowrate 48 ml/min ( . ) continuing with flowrate 6 ml/min ( • ) . combined effect of the mass transfer resistance within the liquid phase and the kinetics of the desorption reaction on the overall elution rate. It must be noticed, on the other hand, that an overall elution time of almost 15 minutes is required to complete the recovery of the adsorbed product, while several hours are necessary when affinity resins are used as stationary phases.
272
Use Of Micro-Porous Aff'mity Membranes For Protein Purification: A Case Study - Sarti
Thus, the use of micro-porous membranes, characterized by an open porous structure, strongly reduces the limitation due to the diffusion of the biomolecules within the liquid phase, allowing shorter processing times for each single stage of the entire separation process. This aspect is one of the main advantages associated with the use of modified membranes as stationary phase in chromatographic processes. Furthermore, higher flow rate can be applied with reduced pressure drops along the column, preserving the mechanical stability of the stationary phase and the biological activity of the immobilized product [19,
20]. The binding capacity has been determined from the amount of recovered product; the value obtained per unit total external area is equal to 0.015 mg/cm 2 or equivalently 1.6 10v mmol/cm 2 considering a molecular weight of 97 kDa; the capacity per unit volume of the membranes is 1.05 mg/ml. This value is in excellent agreement with what obtained by using the same stationary phase to purify MBP-fusion proteins of different molecular weights [19, 20]. Interestingly, the capacity observed is also comparable with the typical values encountered for resins characterized by the same specific interaction towards the MBP domain.
4.2
Equilibrium Isotherm Adsorption experiments of pure protein on amylose affinity membranes, performed in a batch system, offer equilibrium data, reported in terms of the pair c and q* in Fig. 4; the curve corresponding to the interpolation of the equilibrium trend with the Langmuir isotherm is also shown. 0.018 0.016 0.014 0.012 ft,, 0.01 0.008 experimental data
•
0.006
m
Langmuir isotherm
0.004 0.002 0
r
i
i
i
i
r
0
0.05
0.1
0.15
0.2
0.25
0.3
C* (m~/ml)
Figure 4. Experimental equilibrium data interpolated by the Langrnuir isotherm. The data are relative to adsorption step of pure MBP-intein-CBD on amylose membranes in batch system.
273
Use Of Micro-Porous Aff'mity Membranes F o r Protein Purification: A Case Study - Sarti
From the fitting of the experimental data one determines the values of the parameters qm and Ka equal to 0.018 mg/cm2 and 0.023 mg/ml, respectively. Steady state data have been also obtained from the series of adsorption experiments performed in the continuos flow configuration. The steady state points (Cinf~ qinf) are reported in Fig. 5 and compared with the equilibrium Langmuir isotherm. This comparison is presented in order to show the effect of convection on the amount of protein adsorbed at steady state. A flow rate equal to 4 ml/min has been applied, in order to realize a surface velocity comparable with that maintained during the adsorption step of the separation process, performed in the membrane column. 0.018 0.016 0.014
f
-~ 0.012 0.01 0.008 0.006 ~X
0.004 0.002
i
0
0.05
i
0.1 c* (m g/m i)
i
0.15
0.2
Figure 5. Comparison between equilibrium and steady state isotherm in flowing systems. Continuous line is the equilibrium Langmuir isotherm. In each pair, the point in the rightmost position (empty marker) refers to steady state conditions, the leflmost point (gray background) refers to the subsequent equilibrium conditions in batch mode.
As clearly evident, there is always a good agreement between the steady state data and the equilibrium curves, even though a certain scatter of data can be observed in the region of low concentration. The effect of the flux of the liquid solution and consequently of its contact time within the stationary phase is negligible when the concentration in the bulk is large enough to reduce the effect of the concentration gradient between the bulk and the solid interface. It is interesting to point out that the equilibrium adsorption isotherm obtained in batch experiments is exactly the same both for new membranes and for membranes previously used in continuous flow adsorption.
4.3
Adsorption Kinetics The concentration of the protein in the liquid solution has been recorded versus time during adsorption. A common experimental behavior is reported in 274
Use Of Micro-Porous Affmity Membranes For Protein Purification: A Case Study - Sarti
Fig. 6. The experimental trend can be interpolated with a suitable kinetic model in order to determine the value of the kinetic constant of the binding reaction k~. In this work we have used a short times analysis, valid under the hypothesis that in the initial adsorption times the desorption reaction is negligible and the adsorption rate is a function of the feed concentration Co and of the total membrane capacity qm (Equation (7)). 0.25
0.2
o.15
~ o.1
0.05
0
0
50
100
150
200
time (min)
Figure 6. Experimental adsorption curve of pure MBP-intein-CBD on amylose affinity membranes performed in batch system, T=4°C. The values of kl relative to the three different series of adsorption experiments are reported in Fig. 7. Apart from some inevitable scatter due to protein concentration measurements, the results obtained are quite consistent with each other. The values of the kinetic constant of the desorption reaction k2 can be then estimated once the kinetic constant kl and the equilibrium constant K,~ are known. In order to test the consistency of the analysis, the characteristic rate of the forward reaction, kl*Co, is compared with the rate of the backward reaction k2 in Fig.8, at all the feed concentrations inspected. As it is apparent, at low concentration, say below 0.05 mg/ml, the rates of the adsorption and desorption reactions are of the same order of magnitude, while at higher concentrations the rate of the adsorption reaction is indeed much larger. Thus the short times analysis performed is valid only for protein concentrations exceeding the value of 0.05 mg/ml; in that range the scatter in kl values is indeed rather reduced (Fig.7). On the other hand, the low concentration range is less interesting for the adsorption stage.
275
Use O f M i c r o - P o r o u s A t T m i t y M e m b r a n e s F o r P r o t e i n P u r i f i c a t i o n : A C a s e S t u d y - S a r t i
0,025 • Series 1
x Series 2 z~Series 3
X
0,020
.-., 0,015
.~ 0,010
*# x
x~
•
0,005
0,000
0
i
i
i
i
i
0,05
O,1
O,15
0,2
0,25
0,3
C0 (mg/ml)
Figure 7. Estimated values of kl using the short times analysis model. Series 1 ( • ) refers to adsorption of MBP-intein-CBD in batch system; series 2 (*) refers to adsorption performed on continuos flow conditions, and series 3 (A) refers to adsorption performed in batch system using membranes previously used in series 2.
The conditions maintained during adsorption are not favorable for the release of adsorbed material, independently of the concentration of product immobilized on the solid surface. That explains why for the recovery of a specifically adsorbed protein, a competing substrate must be added to the elution buffer and its ionic strength must be enhanced. 1,0E-02
.
.
.
.
.
...-
_~_ 1,0E-03
I_':l" •
P
El
1,0E-04
o k l *cO I [ ] 11,3
1,0E-05 0
0,05
0,1
0,15
0,2
0,25
0,3
c o (mg/rnl)
Figure 8. Comparison between the characteristic rates of the forward and backward reactions during adsorption of MBP-intein-CBD on amylose affinity membranes in batch system.
276
Use Of Micro-Porous AtTmity Membranes For Protein Purification: A Case Study - Sarti
The estimation of kl allows, also, the determination of the characteristic time of the binding reaction tR defined from the following relation:
tR =
V/,t
Ak, qmO-O )
where Vi,/A represents the ratio between the volume of the liquid within the pores, V/,t, and the total area of the membrane sheets; 0 represents the surface fraction already occupied by adsorbed molecules [2]. Considering never used stationary phases, 0 is equal to 0; for convenience the relevant parameter values are reported in Table 2. It is interesting to compare the characteristic time of the binding reaction tR with the characteristic time of protein diffusion within the pores tD, and with the residence time within the pores, to [2]. The corresponding values are reported in Table 3 together with the total interaction time, tg, defined as a sum of tD and tR. The apparatuses used in this work operate at a total interaction time larger (membrane column) or much larger (Millipore holder) than the residence time tc; thus the operating conditions are definitely not optimized. Indeed, to ensure efficient capture of the ligand the condition ti<
Table 3. Characteristic times of the membrane based system Millipore Large scale Parameters membrane membrane holder holder 1.8s
tc = vi,, I a t. :V~.t/(Ak, q, 0 ti = tD + tR
0))
7.05 s
41.3 s
45.8 s
45.9 s
47.6 s
47.7 s
277
Use Of Micro-Porous Affmity Membranes For Protein Purification: A Case Study - Sarti
4.4
Elution Kinetics. The elution of adsorbed pure protein is performed by loading maltose buffer. The release of immobilized product has been monitored by measuring the protein concentration in the liquid solution, until the steady state conditions are reached. Good operating conditions lead to concentrated fractions and to a complete removal of the bound protein. The presence of maltose combined with a high salt concentration proves to give appreciable results. As confirmed by the experimental elution curves, Fig. 9 the release of adsorbed material is quite fast, with a pronounced increase of the concentration in the liquid solution within the first few minutes. Furthermore, the protein concentration in the elution buffer is linearly increasing in time during the early elution stages, so that equation (9) can be applied for the determination of the kinetic constant of the desorption reaction. The average value thus obtained is k2 '= 2.4 10.3 s -1. As expected, k2" is indeed much larger than the kinetic constant of the desorption reaction evaluated for the adsorption condition, whose average value is k2 = 2.3 10-4 s-I. 0.018 0.016 0.014
~__~0.012 "~
0.01
.~ 0.008 8
0.006 0.004 0.002
0
10
20
30
40
50
60
time (min)
Figure 9. Experimental elution curve of MBP-intein-CBD from amylose affinity membranes. 5.
CONCLUSIONS
The purification procedure to recover MBP-intein-CBD from cell lysate has been presented, using micro-porous affinity membranes containing amylose as affinity ligand for the MBP domain. The adsorption and elution kinetics have been analysed, and the adsorption equilibrium isotherm has been determined. The use of micro-porous amylose membranes as stationary phase for the chromatographic process allows for the recovery of the pure product, in a rather high concentration. The peak of maximum concentration is observed during the 278
Use Of Micro-Porous Affinity Membranes For Protein Purification: A Case Study - Sarti
first instants of the elution step, indicating the good efficiency of the elution buffer applied for the removal of the adsorbed material, as well as the effectiveness of convective flow through the pores in reducing mass transport resistances. One of the main advantages offered by affinity membranes, in comparison with affinity resins, is associated to the convective flow through the membrane pores of the solutions used in the subsequent process stages. That greatly enhances mass transport kinetics and results in a strong reduction of the overall processing time for the separation cycle. Remarkably, that fluid-dynamic effect leads also to a particularly efficient washing step, which is very important for the purity of the final product. The application of higher flow rate during the elution step, on one side results in a lower concentration of the protein in the final product and, on the other side, in a reduction of the total elution time required, so that a proper optimization must be considered. The mechanical stability of the stationary phase is preserved thanks to the reduced pressure drop measured along the column, even at the higher flow rates. Adsorption experiments have been then performed in a batch system using pure protein solutions. The equilibrium data obtained show the typical behavior of the Langmuir isotherm, and the relevant parameters have been calculated. The maximum binding capacity at saturation, resulting from the firing procedure, is qm 0.018 mg/cm2, per unit membrane area, and is 1.26 mg/ml per unit volume of the stationary phase. The value obtained is comparable with the values previously found for other MBP fusion proteins [19-20], when reported on a molar basis. This indicates that the binding mechanism involving the MBP domain is practically the same in the different cases, independently of the other protein domains present in the fusion protein. Steady state conditions obtained from adsorption experiments performed under continuous flow operations have been compared with the adsorption equilibrium curve. Experimental results show that at the flow rate adopted the final amount of immobilized product is independent on the effects of the convective flow at all the concentrations inspected. A better insight on this behavior derives from the comparison of the characteristic times for diffusion in the pores, tD, for the binding reaction, tR, and for convection, tc. For adsorption in both membrane apparatuses, protein diffusion in the liquid phase inside the pores, is by far the fastest process, while the reaction step is order of magnitude the slower. The residence time within the pores is intermediate between tD and tR for the Millipore membrane cell, while is of the same order of magnitude of tR for the membrane column. Therefore, the process kinetics is controlled by the reaction step when the Millipore cell is used; on the contrary, both reaction and convection are important for the membrane column, under the operating conditions investigated. The kinetic analysis performed for the adsorption stage as well as for the =
279
Use Of Micro-Porous AtTmity Membranes For Protein Purification: A Case Study - Sarti
elution step leads to the values of the corresponding kinetic constants. For adsorption the average value of the forward reaction is kl = 0.01 ml/(mg, s) and for the backward reaction is k2 = 2.3 10 -4 s1. The overall adsorption rate is governed by the forward reaction at all feed concentrations above 0.05 mg/ml. At feed concentrations larger than 0.05 mg/ml it is thus reasonable to neglect the unbinding reaction during the loading stage, while the same is not possible in the albeit not practically interesting concentration range below 0.05 mg/ml. The desorption reaction rate, evaluated with the elution buffer, is k2 '= 2.4 10 -3 s1 ; it is obviously much greater than the value of k2 which holds true during the loading stage, due to the presence of maltose as a competing substrate and of a high concentration of NaC1, which enhances the dissociation between protein and ligand. The proper simultaneous presence of maltose and salt allows a rather fast removal of the product, and thus a rather high protein concentration in the recovered fractions, as well as a quite short elution time. 6.
ACKNOWLEDGEMENTS
The authors thankfully acknowledge the contributions of Dr. Cristiana Boi and Professor Georges Belfort, in starting the affinity membrane work, and of Professor Alessandro Hochoeppler for the use of E. Coli strain. Thanks are due also to L. Soia, S. Longhi Calcaterra for their contribution to the experimental work. REFERENCES
[1] [2] [3] [4] [5] [6] [7]
[8]
E. Sada, Engineering Aspects of Bioaffinity Separation. J. Chem. Eng of Japan, 23 (1990) 259. E. Klein, Affinity Membranes, Their Chemistry and Performance in Adsorptive Separation Processes Chapter 10, John Whiley & Sons Inc. New York, 1991. K. G. Briefs, M. R. Kula, Fast Protein Chromatography on Analytical and Preparative Scale using Modified Micro-porous Membranes. Chem. Eng. Sci., 47 (1992) 141. O . K . Kellerman, T. Ferenci, Maltose-Binding Protein from E. Coli. Methods Enzymol., 90 (1982) 459. O.K. Kellermann, S. Szmelchamnn, Active Transport of Maltose in Escherichia Coli K12. Eur. J. Biochem., 47 (1974) 139. P-MAL Protein Fusion and Purification System. Expression and Purification of Proteins from Cloned Genes (New England Biolabs Inc.) C. Maina, P.D. Riggs, A. G. Grandea III, B. E. Slatko, L. S. Moran, J. A. Tagliamonte, L. A. McReynolds, C. Guan, An Escherichia Coli Vector to Express and Purify Foreign Proteins by Fusion and Separation from Maltose-Binding Protein. Gene, 74 (1988) 365. C. Guan, P. Li, P. D. Riggs, H. Inouye, H, Vectors that Facilitate the Expression and Purification of Foreign Peptides in Escherichia Coli by Fusion to Maltose-Binding Protein. Gene 67 (1988), 21
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Use Of Micro-Porous Aff'mity Membranes For Protein Purification: A Case Study - Sarti
[9] [10]
[ll]
[12]
[13] [14] [15]
[16] [17]
[18] [19] [20] [21] [22]
[23]
ImpactVM-CN- One step Purification of Recombinant Proteins Using a Self-Cleavable Affinity Tag, New Englan Biolabs Inc. D. W. Wood, V. Derbyshire, W. Wu, M. Chrtrain, M. Belfort, G. Belfort, Optimized Single- Step Affinity Purification with a Self-Cleaving Intein Applied to Human Acidic Fibroblast Growth Factor. Biotech. Prog., 16 (2000) 1055. F. Cattoli, An Affinity Membrane Process for the Purification of a Class of Fusion Proteins Containing the MBP Domain, PhD Thesis, University of Bologna, Italy, April 2000. L. Sundberg, J. Porath, Preparation of Adsorbent for Biospecific Affinity Chromatography. I. Attachment of Group-Containing Ligands to Insoluble Polymers by Means of Bifunctional Oxiranes. J. of Chromatogr. 90 (1974) 87-98 P. Luby, L. Kuniak, Crosslinking statistics 2a) Relative Reactivities of Amylose Hydroxyl Groups. Makromol. Chem., 180 (1979) 2213. Sodium Borohydride and Potassium Borohydride; a manual of techniques. Metal Hydrides Inc. Beverly MA (USA), 1958. M. M. Bradford, A Rapid and Sensitive Method for the Quantitation of Microgram Quantities of Protein Utilising the Principle of Protein-Dye Binding. Anal. Biochem., 72 (1976) 248. J. J. Sedmak, S. E. Grossberg, A Rapid, Sensitive and Versatile Assay for Protein Using Coomassie Brilliant Blue G-250. Anal. Biochem, 79 (1977) 544. S. Chong, F. B. Mersha, D. G. Comb, M. E. Scott, D. Landry, L. M. Vence, F. B. Perler, J. Benner, R. B. Kucera, C. A. Hirvonen, J. J. Pelletier, H. Paulus, M. Q. Xu, Single-Column Purification of Free Recombinant Proteins Using a Self-Cleavable Affinity Tag Derived from a Protein Splicing Element. Gene, 192 (1997) 271. F. Cattoli, C. Boi, G. C. Sarti, Procedimento, Apparato e Membrane di Affinifft per la Purificazione di Biomolecole. Italian Patent N ° BO2001A00058, 2001. F. Cattoli, G. C. Sarti, Purification of MBP-13 Galactosidase and MBP-Rubredoxin Through Affinity Membrane Separation. Sep. Sci. Technol., 37(7) (2002) 1. F. Cattoli, G. C. Sarti, Separation of MBP Fusion Protein Through Affinity Membranes. Biotech. Progres. 18 (2002) 94. S. Szmelchmann, M. Schwartz, J. Silhavy, W. Boos, Maltose Transport in Escherichia Coli K12. Eur. J. Biochem., 71 (1976) 167. T. J. Silhavy, S. Szmelcmann, W. Boos, M. Schwartz, On the Significance of the Retention of Ligand by Protein. Proc. Nat. Acad. Sci. USA, 72 (6) (1975) 2120. M. E. Young, P. A. Carroad, R. L. Bell, Estimation of Diffusion Coefficients of Proteins. Biotech. & Bioeng. 23 (1980) 947.
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New Insights into Membrane Science and Technology: Polymeric and Biofunctional Membranes D. Bhattacharyya and D.A. Butterfield (Editors) @ 2003 Elsevier Science B.V. All rights reserved.
Chapter 14
Economic production of biopharmaceuticals by high-speed membrane adsorbers W. Demmer, S. Fischer-Friihholz, D. Nuflbaumer and D. Meizner
Sartorius AG Grttingen, Germany 1.
INTRODUCTION
Adsorptive membranes are increasingly used for the isolation and purification of biomolecules, a recent review contains 104 references [1]. Sartobind membranes (Sartorius), which are fiat sheet membranes with a methacrylic polymer grafted onto a microporous support and bearing ionic groups, have found numerous applications at different laboratories [2-13]. For the first generation of this product line [14, 15] the support was a 0.45 ~tm nominal pore size Nylon membrane, which was later substituted by a crosslinked regenerated cellulose membrane of 3 ~tm nominal pore size with a high stability against chemical and enzymatic attack and an extremely low nonspecific adsorption [ 16,17]. This second generation of membrane adsorbers was first used for circular, stacked membrane adsorber modules of up to 0.5 m 2 membrane area [ 18] which were successfully used for semitechnical separations [ 19]. For a further scale up a cylindrical module geometry proved to be superior. Based on this concept a unitized system of modules and hardware was developed, which makes the design of large scale plants with membrane areas up to 100 m 2 and more feasible
[20]. Raising the separation efficiency by maximizing mass transfer is the basic idea when using modified microporous membranes as the stationary matrix in liquid chromatography. Membranes can be converted into efficient adsorbers by attaching functional groups to the inner surface of synthetic microporous membranes. Affinity adsorption, ion exchange or immobilized metal affinity chromatography can be obtained by these membranes. Commercially available are membrane ion exchangers of strong acidic (sulfonic acid), strongly basic (quartemary ammonium), weakly acid (carboxylic acid), and weakly basic (diethylamine) types. A chelating membrane based on the iminodiacetate (IDA) group is applicable for IMAC. The membranes are available in products for lab and process scale. That means data can be generated with lab scale units and used for scale up considerations. For process applications the modules and systems can be adopted to the special needs of the specific separation process to achieve optimal conditions.
283
Economic Production Of Biopharmaceuticals By High-Speed Membrane Adsorbers - Melzner
The use of membrane adsorption techniques offers in relation to conventional column chromatographic procedure some advantages which lead to a better process performance, such as: >' lower manufacturing costs non diffusion controlled exchange kinetics so that higher fluxes are possible easier handling in various module forms >" easier upscaling )~ separation of large biomolecules like endotoxins and virus Especially the capability of binding of large molecules is an advantage in comparison to gel chromatography. Typical gel media have pores in the range of about 30 nm in comparison to 3 - 5 ~tm in membrane adsorbers. Therefore membrane adsorbers can be used for molecules where gel media are not working because of size exclusion effects.
Figure 1: Comparison of a Membrane Adsorber membrane Sartobind Q with Q-Sepharose FF The larger pores of the membrane compared to the beads are clearly visible allowing access of large biomolecules like virus particles to the ion exchange sites located on the inside of the pores.
284
Economic Production Of Biopharmaceuticals By High-Speed Membrane Adsorbers - Melzner
Figure 1 shows a typical 3 ~tm microporous membrane structure in contrast to a commercially available beaded chromatographic material i.e. Sepharose Fast Flow Q-Type. 2.
MODULES AND PROCESS DESIGN
For production and large-scale application the Sartobind ® Factor-Two Family of membrane adsorber modules has been developed. The modules consist of a membrane adsorber reeled up like a paper roll to form a cylindrical module sealed at both ends with POM (polyoxymethylene) caps. For scaling up, the modules have areas between 0.12 m2and 8 m 2. The different module heights can be purchased with 15, 30 or 60 layers of membrane. In combination with the different heights a variety of 15 large-scale modules are available. Since the direction of flow is from inside to the outside of the Membrane Adsorber cylinder, a solid core of the appropriate size is inserted into the module to keep hold-up volume as small as possible. The solid POM cores are as well available in lengths of 3, 6, 12, 25 and 50 cm and the thickness varies with the number of membrane layers used. The module is inserted in a specific housing, which consists of a top and base plate, the housing tube and the solid core. For operating the system the unit is filled first with starting buffer. The feed solution enters the unit at the top (see Fig 2) as indicated by the red arrow.
Figure 2: Schematic cross section through a Factor Two Family module and appropriate housing. The direction of the flow is indicated by the arrows. For details see text.
285
Economic Production Of Biopharmaceuticals By High-Speed Membrane Adsorbers - Melzner
The central cylindrical core distributes the fluid to the inside of the module. The flow is directed from the inner channel radial through the module to the outer channel (see arrows in Fig.2). The permeate leaves the housing at the bottom plate (blue arrow). For large-scale protein isolation the adsorber modules of different sizes can be combined to achieve a desired yield and productivity. The modules offer unprecedented high flow rates and short cycle times in the range of minutes. For process plants, parallel running modular units can be combined with modules in series. A system for the extraction of hemoglobin from bovine blood is shown in Fig 3. Bovine blood as starting material was used as a model system because it is relative safe and easy accessible in large quantities and is representative for numerous biological substances.
ixt ens i on Co re
2 x S-8OK-60-fiO
S-40K-30-50
Set ioi Connectors
S-10K-15-25
Figure 3" Cross section of 21 m2 3 stage pilot plant. Not true to scale (reduced in length to 1/5). 1, venting valves; 2, pressurized air connectors for the actuation of the pressure plates. All components are made from AISI 316L (1.4435) stainless steel, the pressure plate is made from POM.
286
Economic Production Of Biopharmaceuticals By High-Speed Membrane Adsorbers - Melzner
The details of the experiments are recently published [21 ][23] so we will focus here on the special features of parallel and serial connection of modules which makes the system so versatile and attractive. For small multistage plants a serial connector is available, which can be mounted between the housings and leads the flow through from the outer channel of the preceding (upper) stage to the inner channel of the following one. A 21 m 2 three- stage pilot plant based on this concept was assembled in this way (Fig.3) and has been operated successfully. The module equipment is listed in Table 1. Table 1" Module equipment (sulfonic acid type) of the pilot plant used for break through curve experiments.
Stage module type
# of modules
membrane
in Fig. 3
% of total
a r e a [m 2]
area
1
S-80K-60-50
146 + 147
16
76
2
S-40K-30-50
150
4
19
3
S-10K-15-25
152
1
5.0
3.
BREAK THROUGH CURVES
Break through curves for purified hemoglobin were measured at the individual modules (Fig. 4) in separate housings.
2.0
............................
#152 S 10K-15-25
1.8
E ,4 1.2
#14;
1.0 ~e.O.8
#146 S 801<-60-50
= 0.6 0.4 0.2 0.0
" 20
40
t
I
q
60
80
100
120
140
160
180
Protein applied [grams]
Figure 4" Break through curves for the individual modules used in the 3 stage pilot plant shown in Fig. 3. The concentration of hemoglobin in the feed was 1.0 g/L.
287
Economic Production O f Biopharmaceuticals By High-Speed Membrane Adsorbers - Melzner
The flow rate was 1 L/min for the smaller modules # 152 and 150 and 3.5 L/min for the two larger modules # 146 and 147, respectively. Furthermore the break through was determined at the three stage pilot plant which is shown in Fig. 3 and the module set which is given in Table 1 as well as with only the first stage. The feed solution was prepared by dilution of a stock solution of purified hemoglobin by the addition of 5 mM phosphate buffer pH 6.0 to the concentrations indicated in the legends to the figures. The flow rate was 3.5 L/min at a differential pressure of 54 KPa (8 psi) at ambient temperature. The resulting break through curves of the first stage (modules #146 and #147 in parallel connection) and three stage plant are shown in Figure 5 A and B, respectively. The curves for the individual modules in Fig. 4 show a rather steep break through of protein for the two smaller ones and a flatter one for the two larger ones. This is to be expected because the smaller modules adsorb less protein. The dynamic capacity at 5% break through for the individual four modules sums up to roughly 100 grams of Hb. Comparison of the break through curves of the first stage alone (Fig. 5 graph A) with that of the complete three stage plant (Fig. 5 graph B) shows, that a only moderate increase in membrane area (from 16 to 21 m 2, that is 38 %) which is added to the 1. stage results in a tremendous increase of the dynamic protein binding capacity (from 80 g to 140 g, that is 75 %). These results show clearly, that with the Sartobind multistage concept described above the module area and geometry is easily adopted to the specific separation problem. A: Break through curve of the 1. stage (16 m 2 membrane area). The concentration of hemoglobin in the feed was 1.0 ~L. B: Break through curve of three stage pilot plant (16 - 4 - 1 m 2 membrane areas). The concentration of hemoglobin in the feed was 0.86 g/L. 2.5
E c
2
o co
•,., 1.5 c 0
=O.
1
L 0
,', 0.5 <
0
50
100
150
200
Protein applied [grams]
Figure 5: The effect of multistage connection of adsorbers on break through curves, determined with bovine hemoglobin. For details see text.
288
Economic Production Of Biopharmaceuticals By High-Speed Membrane Adsorbers - Melzner
Connecting modules in parallel increases capacity and flux at the sacrifice of performance. On the level of a pilot plant of 21 m 2 it has been shown, that this can be overcome by combinations of parallel and serial connections. By carefully balancing these features multistage plants can be adopted to the specific demands of a given application. Scale up: For the construction of the lab scale and process scale modules the same membrane with only different amount of layers is used. This facilitates the establishing of up scale concepts, i.e., the dimensioning of larger units with the same type of membrane [22]. Furthermore the validation of such processes gets easier because all process modules are tested and fully validated at the factory. 4.
APPLICATIONS
4.1
DNA removal with Membrane Adsorbers
During the downstream processing of proteins derived from biological sources a separate ion-exchange chromatography step has to be accomplished to remove contaminants like residual cellular DNA and its fragments. Due to the small concentration of the contaminant, small bed volumes would be sufficient to remove the impurity. Because of the low volumetric flow rates of small gel columns, the bed volume must be drastically oversized to achieve reasonable process times. As a consequence, more product may be lost due to nonspecific adsorption to the gel and larger hardware has to be used. Such large columns cannot be run economically in a single use mode and much effort is needed to validate the regeneration of the gel material.
200
187
150 DNA (Pg mg-1 100protein) 50-
<4 DNA concentration after affinity step
Aider DNA removal step with Sartobind Q
Figure 6: Clearance of DNA out of 12.500 liter batches to below detection limit with module Q-20K-15-50. The values represent the average from three batches.
289
Economic Production Of Biopharmaceuticals By High-Speed Membrane Adsorbers- Melzner
Membrane Adsorberss provide in contrast high flux and small adsorptive bed volumes. This system is designed for the fast removal of impurities from large feedstreams. Although regeneration is easily possible by standard procedures, disposal of the modules after use can contribute to overall process economics by avoidance of comprehensive validation of the regeneration step. The Factor-Two Family system with the Sartobind Q module has successfully been employed for DNA removal in the industrial production of monoclonal antibodies. The binding capacity for DNA was found to be 156 ~tg DNA per cm 2 which represents 5.6 mg per mL membrane volume. The time needed for the process could be reduced 23-fold in comparison to a gel column. Fig. 6 displays the average of three 12,500 liter Protein A purified monoclonal antibody batches before and after use of module Q -20K-15-50. DNA was cleared below the detection limit. Regeneration and validation issues were eliminated by single use of the modules [24]. 5.
VIRUS PURIFICATION
5.1.
Rapid Purification of Alphaherpesviruses [8]
The purification of two alpha herpesviruses, pseudorabies virus (PrV) and bovine herpesvirus 1 on cation exchange membrane was achieved. Cell culture supematants were passed over strongly acidic membranes and adsorbed virions were eluted with a potassium chloride-containing buffer. Over 85 % of the virus was eluted within a single fraction and specific infectivity of the resulting virus preparation was over 10-fold higher than that of sucrose gradient-purified virions. Cation exchange was also used for purification of pseudorabies virus (PrV) mutants deleted in several glycoproteins which grow in cell culture to titers 10- to 100-fold lower than those obtained by wildtype PrV. For PrV, the presence of non-essential glycoprotein gC, which mediates interaction of virions with cell surface heparin sulfate during attachment, was crucial for the successful purification by cation exchange. For purification, the infected cell culture supematant was diluted 1:2 with MES (20 mM 2-morpholinoethane sulfonic acid, pH 6.2) and applied with a 50mL syringe to a S 100 Sartobind ready-to-use unit equilibrated in MES (Fig. 7) PrV-Ka (strain Kaplan), BHV-1 (bovine herpesvirus 1), PrV-gD- Pass (lacks protein gD) and PrV-gC(lacks the non-essential glyocprotein gC which mediates the interaction of PrV with cell surface heparin sulfate) were eluted mainly with 400 mM KC1 in MES buffer. The binding capacity achieved with the Sartobind ® S 100 unit (Fig. 8) for PrV-Ka was determined. Cell culture supematants from three culture flasks of infected porcine kidney (PK) cells were passed sequentially over the S100
290
Economic Production Of Biopharmaceuticals By High-Speed Membrane Adsorbers - Melzner
membrane. Infectivity in the unbound fraction was determined by titration. Results are summarized in Table 2.
Figure: 7 Application of a virus containing feedstream to a S 100 (strongly acidic) membrane adsorber unit. The liquid is passed through the unit by gentle application of pressure with a disposable syringe. Washing and elution steps are performed in a similar manner.
Specific infec~ity of PrV-Ka = 8E+11
~ 6E+11
I mpP~vP'rote~nt
4E+11
~ 2E+11 ca. 0
Sartobind $100 Sucrose purified gradientpudfled
Figure. 8: Specific infectivity of pseudorabies virus strain Ka adsorbers and sucrose density gradients.
291
purified by membrane
Economic Production Of Biopharmaceuticals By High-Speed Membrane Adsorbers - Melzner
Table 2: Determination of binding capacity of S 100 units for PrV-Ka particles. Specific infectivity in pfu: plaque forming units
pfu~applied
pfu in flowthrough
% pfu bound toSl00:
Flask 1
1.7 x 10 l°
4,0 x 102
1O0
Flask 2
1.7 x 101°
7.Ox109
59
Flask 3
1.7 x 1010
L 6 x l O 1°
6
~
Only minimal amounts of infectious material ware detected in the flow through from the first flask. After loading supematant from the second flask, about half of the material did not bind to the S100 and only 6 % of the third flask. The total binding capacity was approximately 2.5 x 108 pfu/cm 2 membrane. The preparations were investigated regarding polyacrylamide gel electrophoresis and Western blots, transmission electron microscopy and specific infectivity. 6.
RESULTS -First report about the ion exchange purification of herpesviruses -Virus yields and purity are high -Conditions applied are mild: Specific infectivity of purified virion preparations is 10 times higher compared to sucrose purified virions -Rapid procedure which can be completed within 3 h from starting with infected cell culture supematant can be easily performed under sterile conditions if necessary. -High flow rates combined with low back pressure and ease of handling.
7.
REMOVAL OF ENDOTOXINS
Endotoxins (lipopolysaccharides from the membrane of gram negative bacteria) make up the majority of pyrogens which must be removed from pharmaceutical products, biologicals for injection and some media for tissue cultures. Endotoxins in solution range from molecular weights of 10,000 to 1,000,000. However, since proteins are in the same molecular weight range as endotoxins, ultrafiltration cannot be used to separate the endotoxins from the target protein. Endotoxins have a negative charge in most solutions. Therefore, ion exchange resins are the most common depyrogenation method for proteins.
292
Economic Production Of Biopharmaceuticals By High-Speed Membrane Adsorbers - Melzner
Using the strongly basic ion exchanger type Q and a buffer pH lower than the pI of the protein, endotoxin will be bound and protein will pass through the membrane. Endotoxin levels were determined in these tests with a gel clot method of BioWhittaker using serial dilutions of samples when appropriate. Sensitivity of the Limulus Amoebocyte Lysate (LAL) method was 0.06 EU/ml (1 EU/100 Pg). 8.
D E P Y R O G E N A T I O N OF W A T E R
Sartobind: Q15 Sample: 500 ml water for injection was spiked with 1 ~tg/mL (10,000 EU/mL) endotoxin from E. coli (055:B5). Procedure: For depyrogenation before use 20 mL of 1 N NaOH was filtered through the Q 15 and the membrane was kept in contact with the alkali for one hour. The Q 15 was then rinsed with 100 mL pyrogen-free water. An LAL control test on the last drops of the rinse solution showed that the unit was pyrogen-free. The sample was applied to the Q 15 at a flow rate of 50 mL/min and after each run the unit was depyrogenated as indicated above. Three duplicate tests were run with the same unit. Results: In each run there was a reduction of endotoxin from 10,000 EU/mL to below 0.06 EU/mL (detection limit). 9.
D E P Y R O G E N A T I O N OF PROTEIN
Most monoclonal antibodies (mAbs) have a pl in a range of 7-8. Bind the endotoxin using strongly basic anion exchanger and allow the mAb to pass through. Sartobind: Q 15. Sample: 200 mL of 85 - 95 % partially pure mAb (pl 7.4) with a protein concentration of 500 ~tg/mL in 20 mM Tris-HC1, pH 7.5 was spiked with endotoxin from E. coli (055:B5) to a concentration of 100 EU/mL. Depyrogenation and rinsing was done as described above. 200 mL of the sample was passed through a Q 15 unit at a flow rate of 20 mL/min using a peristaltic pump and the filtrate was assayed for endotoxin by LAL. After washing 2 times with 20 mL the protein was eluted with 6 mL of 1 M KC1 and the eluate was assayed for endotoxin by LAL Protein concentration was determined using BCA (Pierce). The test was performed in triplicate. Results: In each run there was a reduction of endotoxin from 100 EU/mL to below 0.06 EU/mL (detection limit). This corresponds to a log reduction >3.2.
293
Economic Production Of Biopharmaceutieals By High-Speed Membrane Adsorbers - Melzner
10.
CONCLUSIONS
The demonstrated membrane adsorber systems represent versatile and economic tools for applications in the biopharmaceutical field and show the possibilities to work in areas where conventional purification systems fail. This is especially important in new or growing fields in the pharmaceutical and biotechnological industry, i.e., isolation of viral vectors or even particles, antibodies or nucleic acid molecules. Additional benefits can be defined as easy set up and handling of a fully validated system which can be used in a "single shot" mode reducing validation cost to a minimum. REFERENCES
[1] [2] [3] [4] [5] [6] [7] [8] [9] [10] [11] [12] [13] [14] [15] [16] [17] [18] [19] [20] [21] [22] [23]
C. Charcosset, J. Chem. Technol. Biotechnol. 71 (1998)95-110. I. Adisaputro, Y. Wu, M. Etzel, J. Liq. Chrom. and Rel. Tech. 19(9) (1996) 1437-50. M. Belanich, B. Cummings, D. Grob, J. Klein, A. O'Connor, D. Yarosh, Pharmaceutical Technology 20(3), (1996) 142-150. S. Broverman, G. Prestwich, BioTechniques 19, (1995) 874-875. B. Champluvier, R.-M. Kula, Bioseparation 2, (1992) 343-351. W. Demmer, H. Hoed, A. Weiss, E. Wuenn, D. Nussbaumer, in: Proceedings of the 5th European Congress on Biotechnology, Vol 2 pp 766-769. C. Christiansen et al, (eds.) Munksgaard, Copenhagen, July (1990). K. Gebauer, J. Thoemmes, R.-M. Kula, Biotech. and Bioengin. 53(3), (1997) 181-189. A. Karger, B. Bettin, H. Granzow, T. Mettenleiter, Journal of Virological Methods 70, (1998) 219-224. D. Luetkemeyer, M. Bretschneider, H. Buentemeyer, J. Lehmann, J. Chromatography 639, (1993) 57-66. O. Reif, R. Freitag, J. Chromatography A 654, (1993) 29-41. O. Reif, R. Freitag, Bioseparations 4, (1994) 369-81. X. Santarelli, F. Domergue, G. Clofent-Sanchez, M. Dabadie, R. Grissely, C. Cassagne, J. Chromatography B 706, (1998) 13-22. T. Sellati, M. Burns, M. Ficazzola, M. Furie, Infection and Immunity 63(11), (1995) 4439-4447. Sartorius AG U.S. Patent 5,215,692 (1993). Sartorius AG U.S. Patent 5,556,708 (1996). Sartorius AG U.S. Patent 5,739,316 (1998). Sartorius AG U.S. Patent 5,547,575 (1996). Sartorius AG U.S. Patent 5,618,418 (1997). W. Wang, S. Lei, H. Monbouquette, W. McGregor, BioPharm. 8(5) (1995) 52-59. Sartorius AG International Application WO 98/41300 (1998). W. Demmer, D. Nussbaumer, presented at BIOEUROPE '98, Bioseparation and Bioprocessing of Biomolecules, Sept. 1998 Cambridge UK. Sartorius AG International Application WO 98/41301 (1998). W. Demmer, D. Nussbaumer, J. Chromatography A, 852 No. (1999) 73-81.
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[24]
J. K. Walter, in: Bioseparation and Bioprocessing; G. Subramanian, (Ed.), Processing, Quality and Characterization, Economics, Safety and Hygiene, Wiley VCH, 1998, vol. II, pp. 447-460.
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Functionalized Membranes for Separations and Reactions
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New Insights into Membrane Science and Technology: Polymeric and Biofunctional Membranes D. Bhattacharyya and D.A. Butterfield (Editors) © 2003 Published by Elsevier Science B.V.
Chapter 15
Polymer grafted membranes S.M.C. Ritchie*
Department of Chemical Engineering, University of Alabama, Tuscaloosa, AL, 35487-0203 * Corresponding author, Phone: 205-348-2712, Fax: 205-348-7558,e-mail: [email protected] 1.
INTRODUCTION
The modification of porous microfiltration and ultrafiltration membranes for control of intrinsic properties is a field of considerable interest. Although these modifications have in general been accomplished through alteration of membrane surface chemistries, considerable interest has also been shown toward the incorporation of polymeric groups and structures into the membrane pores. In the case of dense materials, the porous structure of the membrane behaves as a stabilizing support for grafted polymers that may be too fragile in a homogeneous setting creating a more stable composite material. For example, although hydrogels are very useful for drug release applications [ 1], their utility in industrial applications is limited due to their inability to resist mechanical and osmotic forces [2]. By immobilization of the hydrogel in a microfiltration membrane, mechanical stability is added without significant additional resistance to transport processes in the base hydrogel. Incorporating polymeric groups to form dense and porous polymer grafted membranes can be accomplished either by polymerization in-situ or by grafting of large chain polymers. In either case, the addition of polymeric groups to a membrane can significantly alter properties of the membrane, including: permeability, molecular weight cut-off (MWCO), fouling resistance, separation capabilities, biocompatibility, conductivity, responsiveness to environment, and sorption capacity. The purpose of this chapter will be to describe the membranes used for this purpose, the techniques utilized for modification, the behavior of the modified membranes including some basic models for transport, and finally to give some example applications where this technology is currently used, as well as some future possible applications. 2.
MEMBRANES
The field of membranes is extremely broad, and the applicability of different types of membranes for polymer grafting is equally diverse. Polymeric 299
Polymer Grafted Membranes - Ritchie
membranes of nearly every variety have been utilized in the literature as substrates for polymer addition.. The membranes are generally microfiltration (0.02- 10 ~m) or ultrafiltration (1,000- 100,000 MWCO) varieties and may be configured as flat sheets or hollow fibers. Simple non-woven materials may also be employed. Many different types of inorganic membranes (tubular and flat sheet varieties) have also been utilized, including some inorganic/organic hybrids. In general, polymers may be functionalized in any matrix where the porous structure provides sufficient interaction between the membrane and the grafted polymer in the integrated material.
2.1
Polymeric Membranes Membrane materials utilized for polymer grafting may be of nearly any variety. There are examples in the literature for polymer grafting to fluoropolymers [2-11], polyolefins [12-27], cellulosics [28-32], polysulfones [33], polycarbonates [34], and polyacrylates [35,36]. A summary of these different materials, as well as chemical structures, is given in Table 1. Notice that these are all fairly stable chemically, since they only behave as substrates for the grafted polymer. Thus, only a select few contain reactive groups, including the cellulosics and PAN. In most cases they serve as inert backbone polymers for the more fragile grafted polymer, and thus reactivity is not desired. In addition, the polymers are relatively radiation stable so that material integrity is not overly compromised during pretreatment [37]. The physical properties of the materials add further credence to this philosophy. Since most applications involve interaction with aqueous solutions, the solubility parameters are well away from that of water (47.9 MPal/2). This helps to prevent swelling of the membrane, thus providing a stable constraining matrix for the polymer gel. The polymers are also generally temperature stable. This is particularly true in the case of PES and PC. Although most membrane operations are performed at or near room temperature, emerging applications may benefit from added temperature stability. A summary of these physical properties is given in Table 2. The porous structure of polymeric membranes is conducive to a role as a support structure for polymer grafts. Representative scanning electron micrographs for cellulose acetate and polyethersulfone membranes are shown in Figures 1 and 2. Notice that a network of interconnecting pores is formed during the process of phase-inversion. This network is irregular and there are no straight through pores. This ensures excellent interaction between the membrane substrate and the grafted polymer, such that there is effective transfer of mechanical stability. In addition, there is less chance for the formation of pinhole defects that can nullify separation in dense membranes. Careful control of phase-inversion conditions allows regulation of pore size, porosity, and mesh
300
P o l y m e r Grafted Membranes - Ritchie
size. This network can be utilized as a skeletal support for any subsequently formed hydrogel or polymer network. Table 1" Chemical Structures of Various Membrane Materials for Polymer Grafting Polymer Structure
Fluoropolymers: Poly(tetrafluoroethylene) Poly(vinylidene fluoride)
(PTFE) (PVDF)
-.Ec F2 --CF2 ~n ..~CH2__CF2 ~." n
Polyolefins: Poly(ethylene) Poly(propylene) Cellulosics: Cellulose triacetate
-ECH¢--CH2~ -ECH,--~;H~n CH3
(PE)
n
(PP)
(CTA)
/ Ac
C.Hz
-Eo o~. Ac OH CH=
Cellulose
~oo~..
(C)
OH
Polysulfones: Poly(ether sulfone)
0
(PES)
II
0
Polycarbonates: Polycarbonate
(PC)
Polyacrylates: Poly(acrylonitrile)
(PAN)
CH3
-Ec.,-~. 3-
c.
301
Table 2: Physical Properties of Membrane Materials Utilized for Polymer Grafting (data from [38,39]) Polymer PTFE PVDF PE PP PES CTA C PC PAN
T~~(K) ~5b (MPa1/2) 390 23.2 223-371 12.7 148-275 16-17 170-370 18-19 498 22.9 473 25-28 243-433 18-32 420 383 25.3
Tg = glass transition temperature solubility parameter
51.tin Figure 1: SEM of a cellulose acetate MF membrane.
Figure 2: SEM of a polyethersulfone MF (0.22 pm pore size) membrane filter (Courtesy of Millipore Corporation).
If a more open support is required, non-woven supports (lower half of Figure 3) may be utilized. Since the support has a much coarser structure, local interactions between the support and the grafted polymer would be fewer, and thus there may be less mass transfer resistance. The same would be true for track-etched polycarbonate membranes. Notice in Figure 4 that the pores pass in a linear fashion through the membrane. This would be particularly useful from a fundamental point-of-view, since modeling the flow behavior in these membranes would be easier than comparable (pore size) phase-inversion membranes.
302
P o l y m e r Grafted M e m b r a n e s - Ritchie
100 ~m
Figure 4: SEM of a track-etched (uniform pores of 0.05 - 12 ~tm), Polycarbonate, ISOPORETM membrane filter (Courtesy of Millipore Corporation).
Figure 3: Cross-sectional view of a thin film composite membrane (lower section is the non-woven support).
2.2
Inorganic Membranes The use of inorganic membranes has generally been relegated to systems where superior membrane strength is required. The physical strength of these membranes is unmatched, particularly for higher temperature applications. In addition, inorganic membranes (notably silica) possess excellent acid stability. Surface functionalization of inorganic and organic/inorganic composite membranes has been achieved with negligible loss of mechanical stability through the addition of functional silanes [31,40]. This is in contrast to cellulose acetate membranes (Figure 5), where surface functionalization is achieved via hydrolysis and oxidation. Notice the ragged structure of the membrane compared to the precursor. Mediation of membrane degradation by using milder hydrolysis conditions must be balanced with the potential loss of aldehyde sites used for polymer grafting.
Figure 5: Surface views of a cellulose acetate MF membrane before and after hydrolysis (used with permission, [30]). Inorganic membranes that have been utilized in the literature for polymer grafting include tubular membranes composed of silica [41] or
303
Polymer Grafted Membranes Ritchie -
zirconia [42,43]. Other alternatives include sintered glass [44], sol-gel membranes [45], and inorganic/organic composites [32]. In the latter case, membranes are extruded films of up to 80% silica in polyethylene. The polyethylene is used as a binder, and allows for excellent material flexibility while maintaining acid stability and ease of functionalization [31]. A representative scanning electron micrograph of this latter material is shown in Figure 6.
•
~
~3 ,) ,3 : 3 / :
~
;, ~
~
.5 .3..)
~
~
.
~. c ~ ' , ; :.
Figure 6: Surface view of a silica-polyethylenecomposite MF membrane. 3.
POLYMERIZATION TECHNIQUES
There are many techniques for polymer grafting. Some of these techniques, such as electrochemical and ,/-irradiation, require continuous application of a current or radiation throughout the polymerization process. Others, such as UV-initiation, electron beam, plasma, and silanization, create reactive sites for subsequent formation of polymers in-situ, in general by free radical polymerization. Whole polymers can also be incorporated, either by cross-linking of a polymer solution in the pores, or by specific attachment of polymer chains to the membrane pore surfaces. The basic principles of each technique will be described in detail, as well as the benefits and drawbacks to utilization.
3.1
Electrochemical Polymerization The technique of electrochemical polymerization is based on application of a current in a standard three-electrode cell where a membrane has been affixed to the working electrode [45]. When the current is applied, the polymer forms at the exposed working electrode surface and propagates until the void spaces in the membrane are filled. A schematic of this process is shown in Figure 7. Polymers formed in this manner are commonly polyaniline and polypyrrole
304
Polymer Grafted Membranes Ritchie -
Reference Electrode
1
Counter Electrode
1
Aqueous AnilineElectrolyte Solution
Membrane
Working Electrode
Polyaniline
Figure 7" Schematic of electrochemical polymerization process. (Figure 8). Incorporation of nitrogen on the chain permits quatemization and thus charging of the chain. Counterions may be monoanions or polyelectrolytes for higher or lower electrochemical activity, respectively [46].
H
I
H (a)
(b)
Figure 8: Structures of (a) polyaniline, and (b) polypyrrole. The greatest advantage of electrochemical polymerization is that the resulting polymer, and by association the membrane, is conductive. Polymers made in this fashion are generally more conductive than comparable polymers made chemically [47]. In addition, the response of the conductive polymer is enhanced in the membrane phase [48]. The use of more structured membranes, such as tetraethyl orthosilicate (TEOS) sol-gel films, may be 305
Polymer Gratted Membranes - Ritchie
useful for MEMS applications, for example the production of an array of nanoelectrodes [45]. It should be noted that electropolymerization produces a fairly dense polymer, and thus even monoanion transport is restricted. Consequently, different techniques for polymer grafting are required to make more open structures. 3.2
y-irradiation Another technique for polymer grafting is 7-irradiation. This form of radiation is generally more powerful than UV-radiation, and thus is more able to activate the entire material [18]. During polymerization, a capsule containing monomer and the membrane undergoes exposure to y-irradiation from a Co-60 source. Free radicals formed on the membrane act as initiation sites for polymerization. However, the monomer is also activated in solution, thereby promoting homopolymerization [17]. Since the monomer solution is both outside and inside the membrane (excess monomer solution), pore blockage and transport of additional monomer into the membrane matrix may be an issue. Cross-linking the structure during formation has been found to mediate this effect and create a more homogeneous graft across the membrane thickness [7]. Another disadvantage to this technique is that the source is more expensive and incurs more hazards than UV-sources. However, the extensive use of 7-irradiation for sterilization does present an opportunity for wider usage in polymer grafting. 3.3
UV-Initiation Inter-connected polymer networks (IPNs) are generally formed using radiation-grafting techniques. A crude schematic of an IPN is shown in Figure 9. Notice how the fibers of the support material are interwoven with the grafted polymer network. The network itself is composed of polymer chains that have been stabilized by the introduction of a cross-linking agent in the reaction mixture. Photo-initiators have been utilized in the literature for these reactions, and thus UV-light is required. This technique can be utilized on large areas of membrane, is suitable for continuous processes, and generally activates sites deeper into the membrane than surface activation techniques (e.g., plasma). Polymerization takes place primarily by a free radical mechanism with sites on the membrane material or by activation of photo-initiators. Since the membrane has been thoroughly wetted before initiation, polymerization takes place only in the membrane porous structure [2]. This technique is particularly useful for vinyl-containing monomers. Specific examples of polymers grafted in membranes by UV-initiation and other radiation techniques are shown in Figure 10. It should be noted that these membranes
306
Polymer Grafted Membranes - Ritchie
have been primarily examined in the literature for pH- and ionic strengthresponse, and thus the presence of ionizable groups is imperative.
Crosslinked
Membrane
Network of Grafted Polymer
Support Fiber
Figure 9: Schematic of inter-connected polymer network in a porous support.
Polymer
Radiation Technique
Polystyrene
"Ec"'-~~n
Poly(4-vinylpyridine) Poly(acrylic acid)
~/ UV, y, e-
•
n
UV, y, e-
COOH
Poly(acrylamide)
-ECH=--~H ~n
UV
C~-O
I
NH=
Poly(glycidyl
-Ec.,-~. ~n
methacrylate)
c=o i 0
I
/CH
°:,cl., Figure 1O:Typicalpolymersradiation-graftedinto membranes. 307
e
.
Polymer Grafted Membranes - Ritchie
3.4
Electron Beam
The wide availability of electron beam sources has lead to greater implementation in industry, most notably for curing applications, compared to T-irradiation. Electron beam processing for polymer grafting has generally been used for preirradiation of the membrane. This helps to avoid issues with homopolymerization observed with T-irradiation. Decay of radicals on the membrane surface, however, does necessitate the immediate immersion of the activated material into a monomer solution after pretreatment. Electron beam activation does provide a uniform distribution of active sites throughout the starting material. Although specialized equipment is required, the process is amenable to use in a continuous process. The added capital expense and operating cost for high accelerating potentials (-200 kV) does detract from this technique when compared to UV-initiation. 3.5
Plasma
Electron beam and other radiation sources can alter the intrinsic properties of the membrane, and thus alternatives, such as plasma activation, have been examined. Plasma activation is accomplished by formation of an ionized gas (plasma) at the surface of the membrane. A schematic depicting this operation is shown in Figure 11. The plasma will only activate exposed areas, so it is generally useful more for surface treatment (see Figure 12). However, researchers have utilized it successfully for track-etched membranes, where there is good accessibility to the porous areas of the membrane (recall Figure 4). It should be noted that plasma treatment can result in a wide variety of surface chemistries and chemistries for polymerization. For this reason, it is often difficult to characterize what will occur upon plasma treatment. However, its lower cost compared to Tirradiation and electron beam techniques may make its use in niche areas appropriate. "
Gas Inlet
mY
" ~
~
;,
,,
,~i
: ,|
i,:
: , :~,~i
i~.i
,,",, ,,~ i
MEMBRANE
,:::,
,
:;,
"= V a c u u m
~
RF Power Source
]
+ Figure 11" Plasma treatment system for activation of polymeric membranes. 308
Polymer Grafted Membranes - Ritchie
% % = m =
%
,
..
Figure 12: Radical groups formed by plasma treatment located near the membrane surface, with much less affect on interior. 3.6
In-situ Crosslink
When in-situ polymerization is difficult, for example if the temperature/ pressure required cannot be accomplished in the membrane material, in-situ crosslinking of an existing polymer may be appropriate. In this technique, a solution of the polymer, as well as cross-linking agents, is allowed to fully wet the membrane. Once excess solution is removed, the solution is allowed to gel by the application of heat or UV [27,33]. One example where this is done extensively is with water-soluble polymers, such as polyethylene glycol (PEG) and polyvinyl alcohol (PVA). Either through the use of silanes (PEG) or glutaraldehyde (PVA), very specific molecular weight chains of these polymers may be cross-linked in-situ. The use of precisely controlled molecular weight distributions for the precursors allows for excellent control of gel mesh size [28,29]. This is crucial for control of drug delivery and other end-uses of the modified material. 3.7
Silanization and Polymerization
While silanes can be utilized for cross-linking of PEG, they may also be used to apply initiating groups to inorganic membranes. This is particularly effective for silica membranes, as hydrated surfaces readily react with chloroand alkoxy-silanes. When these silanes also contain functional groups, such as vinyl and amine groups, the membrane surface can be primed for subsequent polymerization (Figure 1 3 ) . Cohen and co-workers have effectively utilized this technique to modify a number of inorganic membranes for use in oil-water emulsion separations. Vinyl-functional silanes permit the subsequent polymerization of vinyl pyrrolidone, producing a layer of 309
Polymer Grafted Membranes - Ritchie
cH3
OH OH OH
+
C H 3 0 - Si - (CH2) 3 - O C H 2 - CH - CH 2
i
\o /
CH30
S i - (CH2) 3 O C H 2- C H - CH 2
"l-
H2N- R
Si - (CH2) 3 - O C H 2 - CH - CH 2 - N H - R
I OH Figure 13: Preparation of silica composite membranes by silane derivatization followed by polyamino acid functionalization. covalently attached chains. These "f'mgers" extend into the membrane pores, inhibiting the passage of oil droplets [42,43]. Previous studies by researchers in Japan have demonstrated the viability of this technique for attachment of solvent-responsive acrylate polymers [50]. 3.8
Whole Polymer Grafting
When polymerization in-situ is not possible, whole polymers may be grafted. Many of the same techniques for surface modification may be used, but the polymer is then passed in convective flow through the membrane, with single point attachment via a single end group. This technique has been used extensively by Bhattacharyya and co-workers [30-32] for the attachment of amino acid homopolymers. This technique is a natural extension of enzyme immobilization techniques, and methodologies are similar. For example, cellulosics are utilized by surface modification (via hydrolysis and oxidation) to create aldehyde groups. A Schiff-base reaction may then be utilized for attachment of the polyamino acids via the terminal primary amine group. The formed double bond is reduced to form an acid-stable imine bond. In the case of inorganic membranes, introduction of a glycidyl-functional silane permits direct formation of the imine bond. In both cases, regular polymeric structures can be introduced into the structure while preserving the intramolecular
310
Polymer Grafted Membranes - Ritchie
interactions that impart special functionality to the membrane. An example of this is the ability to form helical structures, such that exceptional selectivity of Pb(II) over Cd(II) is observed (nearly irreversible sorption of Pb(II)) [31]. In addition, more selective side groups may be utilized, for example thiol groups (cysteine) which are effective for Hg(II) sorption [32]. One caveat to this technique is that immobilization becomes more complicated as molecular weight increases since the sole end group (used for attachment) may get hidden in a globular conformation. 4.
RESPONSIVE MEMBRANES
In all of the above cases, the overall goal is to impart selectivity or to change the properties of an essentially inert and non-selective membrane. Ideally, this selectivity would be tunable to the needs of the user. For example, permeability controlled by temperature, pH, ionic strength, and other environmental conditions may be especially useful for biomedical applications. The goal of this section will be to detail membranes that have been developed in the literature for controlled responses in relation to environmental conditions.
4.1
pH
Control of membrane properties by precise pH control is the most easily visualized concept. The definite pKa for each functional group incorporated into the membrane structure allows for a very precise break in the permeability of a polymer grafted membrane. For example, the pKa for the carboxylic acid
/
[H ÷] [salt] [Me x÷]
/ Figure 14: Environmentaleffects on immobilizedpolyelectrolytes in a membranepore. groups of polyaspartic acid is around 3.9. This is on a logarithmic scale, and thus small changes in the pH result in massive changes in the fraction of
311
Polymer Grafted Membranes - Ritchie
charged and uncharged groups. As the number of charged groups increases, repulsion along the polymer backbone increases such that the attached polymers extend into the membrane pores. This effectively decreases the pore size of the membrane, thereby decreasing the permeability of the membrane. This is shown schematically in Figure 14. This phenomenon has been observed experimentally by a number of researchers, and a specific example is given in Figure 15 [12]. The presence of a step change in the permeability is critical as this imparts a switching capability to the membrane. It should be noted that changing the charge of the fixed groups would invert the behavior displayed in Figure 15.
12 A
10
=,
ohm
I
~_er
ed
EE
6 i
2 .....
2
.b-
""
n
n
4
6
8
pH Figure 15: Permeability differences with pH for poly(4-vinyl pyridine) grafted membranes (used with permission, [12]).
4.2
Ionic Strength A similar response is observed with changes in ionic strength. In this case we are specifically limiting discussion to monovalent ions that possess essentially no affinity for the fixed charge. Whereas hydrogen ions possess high affinity for the charged groups below the pKa, changes in ionic strength affect the permeability by charge shielding. That is, once a significant number of charged groups in solution surround the fixed charge, it is less able to interact with neighboring fixed charges. Hence, the repulsive forces generated by neighboring charged species are mediated by "masking" of these charges at excessive solution ionic strength. The source of ionic strength may be a buffer or simply an excess of sodium or chloride ions. In either case, an effect 312
Polymer Grafted Membranes - Ritchie
similar to protein salting out occurs with corresponding collapse of extended chains to the pore walls, as shown in Figure 14. A reversible switching ability would be observed in this situation simply by the permeation of pure water.
4.3
Heavy Metals
When heavy metals are present in the solution, a different possible scenario exists. This is due to specific interactions between the metals and the charged species, such that neutralization takes place by either ion exchange or chelation. In either case, repulsion among charged groups on the attached polymer chains is reduced, and permeability of the membrane will generally increase. Since heavy metals can bind quite strongly well beyond the pKa of the charged groups, the potential exists for creating an irreversible change in the membrane. This may particularly be the case when chelation occurs, since this mechanism is the result of multiple interactions between the sorbed metal and the polymer (very stable complex). Since the metals cannot be desorbed at the operating conditions, the membrane response to the environment would at least decrease and may become totally inactive. 4.4
Temperature
The use of temperature responsive membranes is of particular interest in the biomedical field. Since the polymers utilized generally have a shift at or near body temperature, they are ideally suited for implantable devices, for example in drug delivery modules. The mechanism of operation is governed
(a) A
~ LCST
V
F-
UCST,~ (b)
Critical Concentration Figure 16: Phase behavior of thermo-responsive graftedpolymer chains.
313
Polymer Grafted Membranes - Ritehie
by the lower critical solution temperature (LCST) of the attached polymer. Notice in Figure 16 that the normal upper critical solution temperature is present (higher solubility as temperature increases). However, there is also a temperature above which solvation of the polymer chain decreases again. At this temperature, there is a phase transition that results in shrinkage of the polymer. ~In the case of polyisopropylacrylamide, this transition occurs at around 32 °C and results in shrinkage of the polymer above this temperature. The transition is relatively sharp, and experimentally determined pure water permeabilities have changed an order of magnitude with as little as a degree change in temperature [51 ]. 5. MODELS FOR TRANSPORT IN POLYMER GRAFTED MEMBRANES
Although transport in polymer grafted membranes has been altered by changing feed conditions (i.e., pH, ionic strength, etc.), models for transport have been lax on the explicit inclusion of these effects. In some cases, this is justified by cross-linking of the polymer matrix to inhibit permeability changes due to changes in solution conditions. For example, Anderson and co-workers have examined protein transport through charged and uncharged polyacrylamide grafted membranes and observed negligible change in permeability [2,3]. In other cases, models are adequate for one state, but fail upon transition of the membrane properties [12,52]. In this section, the general equations used to model this transport will be presented, as well as some insights into inclusion of transition effects. 5.1
General Equations Three basic equations have been utilized to model transport in polymer filled membranes. Since the membranes are generally still porous, equations for flow in pipes and packed beds have been adapted. First, the membrane permeability, k, can be calculated from bulk properties using Darcy's Law, k= Q~/ .~ap
(1)
where Q is the bulk flow rate,/z is the fluid viscosity, l is the membrane thickness, A is the membrane cross-sectional area, and Ap is the pressure drop. Notice that since the bulk flow rate and the cross-sectional area are utilized, knowledge of membrane pore size is not required. In addition, all of the parameters may be determined quite easily. Unfortunately, there is little
314
Polymer Grafted Membranes Ritchie -
predictive value, and no information regarding the membrane (besides thickness) is utilized. The permeability can be utilized in the second basic equation by combination with a material balance to obtain the molar flux of solute by a modified form of Fick's Law, N, =~Ap
(2)
where ks is the solute permeability (based on the membrane permeability and the permeate solute concentration), and Cf can Cp are the feed and permeate solute concentrations, respectively. Again, however, there is little explicit information about the membrane in the equation. The pore size can be elicited via a third basic equation, the Hagen-Poiseuille equation,
321ult
(3)
Ap= D2
where D is the membrane pore diameter and u is the actual fluid velocity (accounting for porosity and tortuosity). Equations (1-3) do an adequate job of predicting properties in static systems, but do not capture the transient nature of responsive membranes. That is, they cannot predict the changes in membrane permeability, solute flux, and effective membrane pore radius at different conditions. Rather, experimental data must be taken at each set of conditions. In addition, no gel characteristics are inherent in the equations. Some progress toward including these parameters has been achieved via empirical models [2,3] and models based on first principles [52].
5.2
Accounting for Gel Properties The permeability of the polymer filled membranes will change as the membrane pore diameter changes. Since the presence of polymer in the pores is the cause for the change in pore diameter, the membrane permeability should be a function of how much polymer is in the membrane. This quantity may be expressed as the gel volume fraction, ~, in the pore. Anderson and coworkers have found good agreement between measured and calculated membrane permeabilities based on a power law relation as in Eq. (4) [2]. k = 4.35 x 10-'8~-33' (cm 2)
(4)
Figure 17 shows how Eq. (4) fits permeability data for uncharged, negatively charged, and positively charged polymer filled membranes. However, this 315
P o l y m e r Grafted Membranes - Ritchie
100
A
•~
10
-
1
-
**X•
• Neutral Gel
O Qsstfenr;aArYe~mGi;S
Gel
E
E,r, L O IX
× I
0.1
I
'
'
'
I
II
I
I
I
'
'
'
'
'
0.1
0.01
Polymer Gel Fraction
Figure 17: Power law fit of permeability date for polyacrylamide gels (used with permission, [2]). comes with the caveat that the gels are sufficiently cross-linked such that there is no change in the structure over a range of pH and ionic strength. In addition, accurate prediction of the gel volume fraction is critical. In Figure 17, gel volume fractions were determined experimentally, and thus the calculated permeability accurately reflected what was observed by experiment. To truly make the model predictive, however, the gel volume fraction should be calculated as well. Childs and co-workers have attempted to calculate the gel volume fraction from f'trst principles for dynamic systems. That is, the immobilized gel is allowed to change conformation as dictated by electrostatic repulsions, hydrogen bonding, and van der Waal's forces. Mika and Childs have proposed calculation of a correlation length based on polymer persistence length (Lp), contour distance (distance between adjacent charged groups, A), and the concentration of charge-beating polymer segments (c) [52]. The resulting correlation length, ~, is then assumed to be the diameter of a hard sphere (2R) about which exists a solvent film. A packed bed composed of
316
Polymer Grained Membranes Ritchie -
these solvated spheres is then assumed to represent the polymer filled pores of the membrane as shown in Figure 18. Since the solvated spheres take up the
Re~ I~~) % ~S
~R
Figure 18: Happel model for packed bed (one pore) of solvated polymer chains as representative gel structure where R is the polymer hard sphere radius and Re is the radius of the spherical solvent envelope. entire volume of the bed, the polymer gel fraction can be related to the correlation length by the following equation,
I 1 / -~/4 = Lp + 167r~bAC) (4m4c)'/8(Ac) -3/4
(5)
where lb is the Bjerrum length. The permeability is then calculated using the equation developed by Happel [56],
k = 9¢
3 + 2¢ 5/3
(6)
'
Mika and Childs found excellent agreement between calculated and measured permeabilities for charged gels. However, the calculated permeability varied by 4 orders of magnitude for the neutralized gel. This indicates the need for inclusion of solution properties such as pH and ionic strength into the model such that neutralization from interactions with hydrogen ions or charge shielding by excessive ionic strength and the resulting compression of the solvent films can be adequately accounted for in the model. This last area has great potential additional research, particularly for inclusion of temperature effects.
317
PolymerG r a f t e d
6.
M e m b r a n e s - Ritchie
APPLICATIONS
Polymer filled membranes can be either static or dynamic with regards to separations. Static polymer filled membranes are generally dense or sufficiently cross-linked such that they are non-responsive to environmental conditions. They are used in applications where long-term stability is required for continuous processing. Dynamic membranes are designed for systems where periodic changes in the environmental conditions trigger a response in the membrane. In this section, applications for static and dynamic membranes will be examined.
6.1
Static Membrane Applications Although there are a wide variety of separations for which polymer filled membranes are applicable, five primary areas will be addressed in this review. These areas include bioseparations and artificial organs, metal sorption and water treatment, fuel cells, nanofiltration, and gas separations/pervaporation. These various applications will be examined starting from porous membranes to more dense membranes. 6.1.1 Bioseparations and Artificial Organs There are two sub-categories in this area of applications: protein separations and immunoisolation. First, protein separations may be performed with polymer filled membranes in a manner very similar to column separation, for example by electrostatic interactions. Depending on the chemistry of the polymer gel, positively or negatively charged proteins may be separated from each other, with the interacting species being retained on the membrane. For example, bovine serum albumin (BSA) has been separating using polyethylene membranes filled with an acrylic polymer that has been modified with amine groups [23]. Interaction of the negatively charged BSA with the membrane allows for its separation. In a similar fashion, by introduction of a sulfonic acid containing polymer, a positively charged species, such as lysozyme, has been successfully separated [22]. It should be noted that to add selectivity to the membrane, more specific interactions would be required, for example the use of ELISA-type groups. However, in comparison to packed bed separations, the polymer filled membranes do offer the advantage of removing larger proteins in convective flow, which eliminates the diffusive mass transfer resistances encountered with ion exchange resins and nanoporous column packings. More intriguing is the use of these membranes for immunoisolation. In that case, the membrane is used to regulate the passage of nutrients, therapeutic agents, and immunological molecules. The object here is the creation of an implantable device, containing living cells, that permits 318
Polymer Gratted Membranes - Ritchie
production and delivery of a therapeutic agent. Since live cells are implanted, there is a requirement for nutrient transport to the cells and for release of the produced therapeutic agent to the affected region. In addition, immunological agents that would destroy the foreign cells must be kept away. A schematic of this process is shown in Figure 19.
•
Immunological Molecules
Therapeutic Molecules Nutrients
Live
Figure 19: Schematicof implanted device containinglive cells (e.g., islets of Langerhans for insulin production). Barbari and co-workers have developed membranes with a cross-linker gradient across the filled polymer to regulate the molecular weight cut-off of the membranes used in this process [28,29]. Glutaraldehyde was used to cross-link a polyvinyl alcohol gel in cellulose ester membranes. In their work, the representative nutrient molecule passed relatively uninhibited. Therapeutic molecules, represented by a 50 kD protein passed through the membrane, though much more slowly than the nutrient. Very large molecules (> 150 kD) were rejected by the membrane to a greater degree, but only with a selectivity of about 2.5 (compared to 50 kD). Similar results have been observed by Baker and co-workers [33]. In that case, polyethersulfone membranes containing crosslinked polyvinyl alcohol were developed for the immunoisolation of pancreatic cells. In that case, in vivo studies of the membranes showed negligible change in membrane performance after 6 months, thereby showing significant promise for their use in implantable devices containing living cells. 6.1.2 Metal Sorption and Water Treatment Another large application for polymer filled membranes is heavy metal sorption for the treatment of wastewater. Work in this area was begun in Japan by radiation grafting of glycidyl methacrylate containing polymers in 319
Polymer Grafted Membranes - Ritchie
polyolefin membranes. Post-modification of these membranes to get iminodiacetic acid (IDA) groups permits the utilization of chelation mechanisms for sorption of metal ions [21]. The presence of IDA groups is particularly useful since they allow for much greater selectivity over ions such as calcium and magnesium. Additional work where membranes are utilized to mimic the behavior of ion exchange and chelation resins has been performed by Choi and Nho [18]. They polymerized styrene in a membrane, followed by material sulfonation. In that case, the membrane is useful over a wider pH range since the membrane mimics a strong acid ion exchange resin. Further enhancements have including copolymerization to include carboxylic acid and pyridine groups for the production of mixed ion exchange and chelation membranes [19,20]. It should be noted that in all of these cases, the polymer is formed in-situ, thereby limiting the use of more complex polymers that possess improved selectivity or stability. In work done by Bhattacharyya and co-workers [30-32], homopolymers of various amino acids have been immobilized in a variety of organic (cellulosic) and inorganic-organic hybrid (silica-polyethylene) microfiltration membranes. Work with polyglutamic and polyaspartic acids has revealed an additional mechanism of metal sorption besides ion exchange and chelation. This mechanism is counterion condensation, which allows non-specific sorption of metal ions by attraction to a negatively charged electrostatic field caused by close residence of charged groups on the polymer [31 ]. Since this is a regular structure, the spacing between charged groups is regular, and superposition of the electrostatic fields in the membrane pores leads to a large overall field effect. The result is exceptionally high sorption of heavy metals with greater than 1 g heavy metal sorbed per gram of dry membrane sorbent [30]. The use of polyamino acids also permits the use of more exotic functional groups, such as thiol groups that have been shown to be particularly effective for the removal of soft metals, such as mercury [32]. 6.1.3 Fuel Cells The use of membranes containing sulfonic acid groups has also found utility for fuel cell membranes. Currently, the benchmark for new fuel cell membranes is Nation® [59]. One of the critical parameters for their use in fuel cell applications is the presence of water in the membrane. Water is critical to maintain conductivity across the membrane. For Nation® there must be at least 6-7 water molecules for each sulfonic acid group in the membrane [4]. Operating conditions usually involve 24 water molecules for each sulfonic acid group. The use of polystyrene that has been subsequently sulfonated is an excellent technique for raising the number of sulfonic acid groups in the membrane, as well as increasing its hydrophilicity and conductivity [4]. Membranes developed by the incorporation of sulfonated
320
Polymer Grained Membranes - Ritchie
polystyrene have led to the inclusion of around 60 water molecules for each sulfonic acid group [57]. The resulting conductivity is about double that of Nation®. Considering the expense of Nation®, this represents an attractive alternative. One caveat that must be considered is that for application in direct methanol fuel cells, the passage of methanol must not be too high, and the porous structure of these modified membranes need to be considered [58].
6.1.4 Nanofiltration Another use for membranes containing charged polymers is nanofiltration. This is a particularly intriguing application, since a microfiltration membrane ( 0 . 0 2 - 10 ~tm pore size) is used to reject ions. Childs and co-workers [27] have incorporated poly(4-vinyl pyridine) (PVP) into microfiltration membranes and achieved 70% rejection of a NaC1 solution at only 300 kPa (< 3 bar). When membranes containing an acrylic polymer gel are utilized, greater than 90% rejection of Na2SO4 was observed. Higher rejection of divalent cationic species should also be observed with the PVP membranes due to their positive charge, making them attractive for water softening applications. It should be noted that since these membranes are fairly open structures, excessively large fluxes will lead to loss of solute rejection. In addition, increases in ionic strength, as with regular nanofiltration membranes, will lead to charge shielding and a subsequent drop in solute rejection. 6.1.5 Gas Separation and Pervaporation In this last application of static membranes, alteration of solute solubility is obtained by changing the membrane chemistry. These membranes are also much more dense than previous examples as pinhole defects ruin membrane selectivity. One example in the area of gas separations is the recovery of CO2 from flue gas streams. Mechanical and thermal stability is provided by the inert membrane, in this case a polyacrylonitrile membrane [36]. A polymer that has excellent solubility for CO2, poly(ethylene glycol) acrylate (PEGA), was then incorporated as a gel in the membrane. Since it is technically challenging to make consistent, thin, stable films of PEGA, this is an ideal combination. When these materials were utilized for CO2/N2 separation at 30 °C, a selectivity of 32.4 was observed [36]. Pervaporation membranes are also dense and transport is governed by the solubility and diffusivity of solutes in the membrane material. For example, Ihm and Ihm [10] have incorporated sulfonated polystyrene in polyvinylidene (PVDF) membranes for pervaporation of ethanol/alcohol mixtures. When utilized as part of a membrane reactor, water formed during esterification reactions at sulfonic acid membrane surface sites (not in the 321
Polymer Grafted Membranes - Ritehie
porous structure) may be subsequently removed from the reacting phase by transport through the membrane. This is illustrated schematically in Figure 20.
Ethyl Acetate
Water
Ben:ene -I),
Cyclohexan
+
Benzene
Ethanol
(a)
(b)
Figure 20: Polymer grafted membranes for (a) water extraction in a membrane reactor, and (b) organic/organic separation by pervaporation.
The selectivity of the membrane was not especially high (only around 21), but used in this manner it can help to limit unwanted side reactions and can maximize reaction rates. A more important application for pervaporation may be in organic/organic separations. Classically, membrane separations have been limited in this capacity, as the formation of mainly polymeric membranes is performed in organic solvents. Swelling on the membranes is thus a major concern, in regards to the loss of selectivity, and in the lifetime of the membranes. The use of polymer filled membranes shows great promise in this area, as an inert support can add the required mechanical strength to resist or control swelling of the separating polymer. Kai et al. [25] made hollow fiber and fiat sheet polyethylene membranes containing polymethacrylate. Flat sheet membranes were found to have a lower flux but higher separation factor, while hollow fibers provided a higher flux with less selectivity. This lead to the conclusion that fiat sheet membranes were better equipped to add mechanical strength and thus control the swelling of the filled polymer. Earlier work by the same group for benzene/cyclohexane separations demonstrated a selectivity of 7-8 due to the enhanced solubility of the aromatic in polymethacrylate [26]. This is significant since distillation is especially costly for the separation of this mixture due to low relative volatility (- 1.02) and the formation of an azeotrope.
322
Polymer Grafted Membranes - Ritchie
6.2
Dynamic Membrane Applications The potential of adding dynamic properties is an exciting development of polymer grafted membranes. Switching capabilities provides great promise for the creation of biomimetic structures, "lab-on-a-chip", and other structures for emerging applications. Not surprisingly, biomedical applications are near the forefront of this technology, with some applications in chemical valves and environmentally stimulated drug delivery. For each of these applications, however, very simple responses, such as pore expansion with shear [41], to more complex operations, where membrane permeability is altered by differences in solution pH, ionic strength, and temperature are utilized.
6.2.1 Oil/Water Separations Cohen and co-workers have examined the use of polymer filled membranes for oil/water separations. In their work, poly(vinylpyrrolidone) has been grafted onto ceramic membranes. Originally this work was performed to decrease the pore size and permeability of ceramic membranes [40]. Due to their method of formation (sintering of f'me ceramic powders), the formation of very narrow pore size distribution ceramic membranes is extremely challenging. This work presented a method of post-treatment that could sufficiently decrease the pore size to enable effect oil/water separation. Since the filled polymer was charged, they found that pH, ionic strength, and solvation of the chains could all be utilized to decrease the membrane permeability and effective pore size. Although they have not been utilized specifically as dynamic membranes, this work does indicate the possibility, in this case for ceramic membranes. 6.2.2 Chemical Valves Dynamic membranes have been used more specifically in the literature as chemical valves. A chemical valve, as shown in Figure 14, has been defined as the reversible expansion and contraction of membrane pores in response to environmental conditions of pH, ionic strength, and divalent/trivalent metals. In all cases, when charged groups interact with ions in solution, the end effect is the same: shrinkage of the immobilized chains and an increase in membrane permeability up to three orders of magnitude [53]. Changes in the membrane permeability are then a function of metal concentration, ionic strength, and pH. If the ionic strength of the solution is held constant, then the permeability has been found to change linearly with pH [53]. Thus, a very simple model,
k=a-b(pH)
(7) 323
Polymer Grafted Membranes - Ritchie
is sufficient to predict the membrane permeability. Unfortunately, this relation does not hold if the ionic strength is not held constant, and the behavior will be more reminiscent of Figure 15. That is, there will be a step jump in the permeability above the pKa [34]. Comparisons to biological membranes are apparent, and the biomimetic characteristics of these membranes makes them excellent candidates for uses in the biomedical industry, in nanosized reactors, for "lab-on-a-chip", and other novel structures.
6.2.3 Thermo-responsive Membranes and Drug Delivery Hydrogels have been investigated extensively for drug delivery applications, and their extension to inclusion in membranes should thus not be unexpected. In this way, some negative aspects regarding hydrogels as drug delivery media, most notably issues with stability, are addressed. Hydrogels consisting of N-isopropylacrylamide are well known to have a volume-phase transition at around 32 °C. Early work with this gel in a porous glass membrane found a dramatic shift in the molecular weight cut-off (MWCO) of the membrane of about an order of magnitude [54]. Once a change in the permeability was established, research began examining the transport of drugs. Conventional thermo-responsive drug delivery systems only provide slight or gradual changes in drug transport rates as temperature changes due to increases in diffusivity and solubility. Polymer filled membranes were used in an attempt to make rate changes more dramatic, thus obtaining a switching mechanism. In this case, the transport of salbutamol sulfate, a pulmonary drug, was examined across hydrophilic and hydrophobic membranes containing a thermo-responsive polymer [55]. In their study, there was some change in the diffusion rate across the membrane at different temperatures. However, there was only a distinct difference for cellulose nitrate membranes. Similar work for BSA transport performed by Li and D'Emanuele [44] gave similar results. Some interesting ideas of better controlling the switching point have been put forward by Iwata et al [51]. They have indicated that by copolymerization of a thermo-responsive polymer with a second monomer that the inflection point may be increased or decreased. 7.
CONCLUDING REMARKS
Polymer grafting offers an excellent opportunity for controlling and tuning the properties of membranes. For example, properties such as permeability, separation capability, conductivity, and biocompatibility can be altered vastly, simply by the attachment of different polymer chains in the membrane porous structure. An additional feature is that in some cases, these changes can be dynamic, with stimuli-response to various environmental 324
Polymer Grafted Membranes - Ritchie
conditions, such as pH, ionic strength, metal concentration, and temperature. Membranes are available in a wide range of compositions and configurations, and thus there are many different techniques for grafting the polymers. In most cases this is by in-situ polymerization, but immobilization of large chain polymers to utilized more exotic moieties has also been accomplished. Many researchers have attempted to model the behavior of polymer grafted membranes, and though they have been successful for static membranes, more work is needed to truly account for stimuli explicitly for dynamic membranes. This is particularly true in the case of thermo-responsive membranes. In any case, the many different areas where these membranes can and are being applied, from filtration to biomedical devices, should ensure that this emerging area in polymer and membrane science will be of interest for years to come. REFERENCES
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N.A. Peppas, P. Bures, W. Leobandung, H. Ichikawa, Eur. J. Pharma. Biopharma., 50 (2000) 27. V. Kapur, J.C. Charkoudian, S.B. Kessler, J.L. Anderson, Ind. Eng. Chem. Res. 35 (1996) 3179. V. Kapur, J. Charkoudian, J.L. Anderson, J. Membr. Sci, 131 (1997) 143. D.I. Ostrovskii, L.M. Torell, M. Paronen, S. Hietala, F. Sundholm, Solid State Ionics, 97 (1997) 315. J. Hautojarvi, K. Kontturi, J.H. Nasman, B.L. Svarfvar, P. Viinikka, M. Vuoristo, Ind. Eng. Chem. Res., 35 (1996) 450. R. Teleranta, J.A. Manzanares, K. Kontturi, J. Electroanal. Chem., 464 (1999) 222. B. Gupta, F.N. Buchi, G.G. Scherer, A. Chapiro, J. Membr. Sci., 118 (1996) 231. K. Kontturi, S. Mafe, J.A. Manzanares, B.L. Svarfvar, P. Viinikka, Macromolecules, 29 (1996) 5740. G. Ellinghorst, A. Niemoller, D. Vierkotten, Radiat. Phys. Chem., 22 (1983) 635. C.-D. Ihm, S.-K. Ihm, J. Membr. Sci., 98 (1995) 89. H. Matsuyama, M. Teramoto, H. Sakakura, K. Iwai, J. Membr. Sci., 117 (1996) 251. A.M. Mika, R.F. Childs, J.M. Dickson, J. Membr. Sci., in press 2002. M. Ulbricht, React. Func. Polym., 31 (1996) 165. A.M. Mika, R.F. Childs, J.M. Dickson, J. Membr. Sci., 153 (1999) 45. A.M. Mika, R.F. Childs, J.M. Dickson, B.E. McCarry, D.R. Gagnon, J. Membr. Sci., 108 (1995) 37. N.M. E1-Sawy, A.Z.A. Elassar, Eur. Polym. J., 34 (1998) 1073. Y.C. Nho, J.L. Garnett, P.A. Dworjanyn, J. Polym. Sci.: Part A: Polym. Chem., 31 (1993) 1621. S.-H. Choi, Y.C. Nho, J. App. Polym. Sci., 71 (1999) 2227. E.-S.A. Hegazy, H. Kamal, N. Maziad, A.M. Dessouki, Nuc. Instr.Meth. Phys. Res. B, 151 (1999) 386. E.-S.A. Hegazy, H.A. Abd E1-Rehim, H.A. Shawky, Rad. Phys. Chem., 57 (2000) 85. S. Konishi, K. Saito, S. Fumsaki, T. Sugo, Ind. Eng. Chem. Res., 31 (1992) 2722. M. Kim, M. Sasaki, K. Saito, K. Sugita, T. Sugo, Biotechnol. Prog., 14 (1998) 661. K. Sunaga, M. Kim, K. Saito, K. Sugita, T. Sugo, Chem. Mater., 11 (1999) 1986.
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S. Tsuneda, T. Endo, K. Saito, K. Sugita, K. Horie, T. Yamashita, T. Sugo, Macromolecules, 31 (1998) 366. T. Kai, T. Tsuru, S. Nakao, S. Kimura, J. Membr. Sci., 170 (2000) 61. T. Yamaguchi, S. Nakao, S. Kimura, Macromolecules, 24 (1991) 5522. R.F. Childs, A.M. Mika, A.K. Pandey, C. McCrory, S. Mouton, J.M. Dickson, Sep. Purif. Tech., 22-23 (2001) 507. W.S. Dai, T.A. Barbari, J. Membr. Sci., 171 (2000) 79. W.S. Dai, T.A. Barbari, Biomaterials, 21 (2000) 1363. D. Bhattacharyya, J.A. Hestekin, P. Brushaber, L. Cullen, L.G. Bachas, S.K. Sikdar, J. Membr. Sci., 141 (1998) 121. S.M.C. Ritchie, L.G. Bachas, T. Olin, S.K. Sikdar, D. Bhattacharyya, Langmuir, 15 (1999) 6346. S.M.C. Ritchie, K.E. Kissick, L.G. Bachas, S.K. Sikdar, C. Parikh, D. Bhattacharyya, Env. Sci. Tech., 35 (2001) 3252. A.R. Baker, R.L. Foumier, J.G. Sarver, J.L. Long, P.J. Goldblatt, J.M. Homer, S.H. Selman, Cell Transplant., 6 (1997) 585. Y. Ito, M. Inaba, D.-J. Chung, Y. Imanishi, Macromolecules, 25 (1992) 7313. M. Ulbricht, K. Richau, H. Kamusewitz, Coll. Surf. A: Phys. Eng. Asp., 138 (1998) 353. J.H. Kim, S.Y. Ha, S.Y. Nam, J.W. Rhim, K.H. Baek, Y.M. Lee, J. Membr. Sci., 186 (2001) 97. A. Dawood, K. Miura, Polym. Degrad. Stab., 73 (2001) 347. Encyclopedia of Polymer Science and Engineering, 2n°Edition, Ed.: H.F. Mark, N.M. Bikales, C.G. Overberger, G. Menges, Wiley: New York, 1985. Polymer Handbook, 4th Edition, Ed.: J. Brandup, E.H. Immergut, E.A. Grulke, Wiley: New York, 1999. R.P. Castro, Y. Cohen, H.G. Monbouquette, J. Membr. Sci., 84 (1993) 151. R.P. Castro, H.G. Monbouquette, Y. Cohen, J. Membr. Sci., 179 (2000) 207. R.S. Faibish, Y. Cohen, Coll. Surf. A: Phys. Eng. Asp., 191 (2001) 27. R.S. Faibish, Y. Cohen, J. Membr. Sci., 185 (2001) 129. S.K. Li, A. D'Emanuele, J. Contr. Rel., 75 (2001) 55. M.M. Verghese, K. Ramanathan, S.M. Ashraf, M.N. Kamalasanan, B.D. Malhotra, Chem. Mater., 8 (1996) 822. W.S. Shim, Y.H. Lee, I.-H. Yeo, J.Y. Lee, D.S. Lee, Synth. Met., 104 (1999) 119. A.F. Diaz, K.K. Kanazawa, G.P. Gardini, J. Chem. Soc. Chem. Commun., (1979) p. 635. S. Das Neves, M.-A. De Paoli, Synth. Met., 96 (1998) 49. S. Machi, Radiat. Phys. Chem., 14 (1979) 155. K. Otake, T. Tsuji, M. Konno, S. Saito, J. Chem. Eng. Japan, 21 (1988) 443. H. Iwata, M. Oodate, Y. Uyama, H. Amemiya, Y. Ikada,Y., J. Membr. Sei., 55 (1991) 119. A.M. Mika, R.F. Childs, Ind. Eng. Chem. Res., 40 (2001) 1694. Y. Osada, K. Honda, M. Ohta, J. Membr. Sci., 27 (1986) 327. T. Tsuji, M. Konno, S. Sato, J. Chem. Eng. Japan, 23 (1990) 447. S.-Y. Lin, Y.-Y. Lin, K.-S. Chen, Pharma. Res., 13 (1996) 914. J. Happel, H. Brenner, Low Reynolds Number Hydrodynamics with Special Applications to Particulate Media, Prentice-Hall:Englewood Cliffs, NJ, 1965. S. Holmberg, T. Lehtinen, J. Nasman, D. Ostrovskii, M. Paronen, R. Serimaa, F. Sundholm, G. Sundholm, L. Torell, M. Torkkeli, J. Mater. Chem., 6 (1996) 1309.
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K. Scott, W.M. Taama, P. Argyropoulos, J. Membr. Sci., 171 (2000) 119. H.-G. Haubold, Th. Vad, H. Jungbluth, P. Hiller, Electrochim. Acta, 46 (2001) 1559.
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New Insights into Membrane Science and Technology: Polymeric and Biofunctional Membranes D. Bhattacharyya and D.A. Butterfield (Editors) O 2003 Elsevier Science B.V. All rights reserved.
Chapter 16
Functionalized membranes for tunable separations and toxic metal capture A. M. Hollman and D. Bhattacharyya
Department of Chemical & Materials Engineering, University of Kentucky, Lexington, Kentucky 40506-0046 * Corresponding Author, e-mail: [email protected], Phone: 859-257-2794, Fax: 859-323-1929
1.
INTRODUCTION
The design of sustainable processes for separations, chemical reaction and heavy metal sorption will require the development of systems with reduced energy consumption and minimal environmental impact. Membrane processes have gained wide acceptance due to their unique properties such as their compactness, ease of fabrication, low-cost operation and modular design. For these reasons, membrane processes are finding numerous applications in fields ranging from water treatment to reactor design to advanced bioseparations. Conventional separations involving pressure-driven membrane processes are classified as: reverse osmosis or RO (for very high degree of salt separation including NaC1, and soluble organics); nanofiltration or NF (for separation of divalent salts and organics while permeating monovalent salts such as NaC1); ultrafiltration or UF (for separation of organics above 5000 MW); and microfiltration or MF (for separation of colloidal to micron particles and microorganisms) [1-2]. For UF and MF type membranes the mechanism of separation is based on size exclusion or steric hindrance factors. For more dense NF membranes this mechanism of exclusion is coupled with the electrostatic repulsive forces resulting from the presence of ionizable functionalities incorporated within the membrane matrix. However through appropriate modification (i.e. with charged macro-molecules), the applicability of porous microfiltration membranes can be extended to molecular level separations at applied pressures well below those required by traditional RO and NF. The development of membrane materials containing immobilized biomolecules (i.e. poly(amino acids), proteins, enzymes etc.) provide added
329
Functionalized Membranes For Tunable Separations and Toxic Metal Capture - Bhattaeharyya
opportunities for high performance selective separations and recognition-based chemical reactions. Traditionally, microfiltration (MF) membrane separations have been limited to the filtration of suspended solids, bacteria, viruses, etc. However, the separation capabilities of MF-type membranes (i.e. cellulose, silica, polysulfone) can be enhanced through functionalization with a variety of reagents. Depending on the chemical nature of the attached molecules (such as its chain length, acidic or basic properties, catalytic reactive site, etc.), these types of low resistance membranes could be used in processes like dissolved metal ion (or oxyanions) separation, chlorinated organic detoxification or biocatalysis. The primary focus of this chapter will be the use of microporous support membranes containing immobilized multi-functionalities for applications involving tunable charged-based separations or the development of high capacity sorbent materials. In particular, the fabrication of these novel membrane materials has been achieved by attachment of various polypeptides (MW 2,500-100,000) directly within the pore structure. Both the configurational and electrostatic properties as well as the binding affinity of these polyligands have been exploited to yield highly specific separations involving aqueous systems. 20
TUNABLE SEPARATIONS
Growing interest, in recent years, regarding the chemical modification of synthetic membranes has resulted in the development of materials that are sensitive to changes in their surrounding micro-environment. The properties of these new generation membranes, such as reversible water permeability [3-5], ion-selectivity [6-8] and surface hydrophobicity [9], have been shown to respond to slight variations in temperature [10], ionic strength [11], electric current [12], or pH. These stimuli-responsive membranes could have extended applications in catalysis, recognition-based separations, biomimetic device development, and controlled-drug release. 2.1
Stimuli-Responsive Macromolecules The most common methodology used in preparing an environmentally sensitive material is to alter the properties of a base support through the immobilization of an ionizable polymer. Monomeric repeat structures of some synthetic and biological polymers used in the development of tunable membranes are shown in Figure 1. Polyelectrolytes, comprised of these repeat units, are particularly well suited for modulating chemical signals because their secondary structure is dependent upon the electrostatic interactions between side group constituents. In solution, these intramolecular forces are regulated by the
330
FunctionalizedMembranesFor TunableSeparationsand ToxicMetalCapture- Bhattaeharyya
o II
O I! O II H2c - - CH - C - OH
acrylic acid
H2C-C-C-OH I CH 3
methacrylic acid
O II H 2 C : CH - C - NH 2
H2C - CH -- C -- NH CHCH 3 1 CH 3
acrylamide
N-isopropylacrylamide
NH H NH20 II ~ ~ II H2N - - C - NHCH2CH2CH2 C .... C - O H
CH - CH 2
4-vinylpyridine
O II HO-C-CH
L-arginine
vinylbenzyl chloride
H2N t r i O °~ II 2 CH 2 C - - C - O H
L-glutamic acid
H2N . H O H2NCH2(CH2)2CH2C -- C -- OH
L-lysine
H2N H O II HSCH2C--C-OH
L-cysteine
Figure 1. Repeat unit structures of some polymeric materials used for the development of stimuli-responsive membranes.
properties of the contacting solution, such as pH, ionic strength and temperature. A typical morphological response to variations in pH for a macromolecule containing acidic side groups (i.e. poly(acrylic acid), poly(methacrylic acid), poly(L- glutamic acid), etc.) is given in Figure 2. At low pH, the carboxyl side groups are protonated alleviating the repulsive force between neighboring constituents. This causes the polymer chain to contract, thus increasing its entropy. Conversely at high pH, these carboxylic acid groups are fully dissociated. Under these conditions, the electrostatic energy overcomes the entropic contribution to the free energy resulting in an extension of the polymer chain [13]. Upon immobilization, this response can be utilized to tune the performance characteristics of membrane materials. 2.2
Immobilization Techniques Preparation of these 'intelligent' or 'smart' materials can be achieved through both covalent and non-covalent immobilization techniques. Typically, these chemical modifications involve either in-situ polymerization [14-16] or the grafting of pre-existing macromolecules [7, 17-18]. 2.2.1 Cross-Linked Hydrogels An example of the former case is the incorporation of cross-linked hydrogels
331
Functionalized Membranes
For Tunable Separations and Toxic Metal Capture - Bhattacharyya
100 % 80
.X
60 \ 40
20
%
% ~
Low
pH
~
mm - - -
mm am=
High
Figure 2. Typical morphological response to variations in the pH of the contacting solution for a negatively-charged macromolecule (i.e. poly(L-glutamic acid), polyacrylic acid, etc.). within porous media [5, 15]. Hydrogels are defined by their ability to undergo large volume changes when exposed to an external stimulus. The usefulness of gels as separation devices is dependent upon the ability to maintain their integrity against mechanical stresses. To circumvent problems associated with stability, hydrogels can be enmeshed within the pores of a rigid polymer matrix by cross-linking [15]. In-situ polymerization is carried out by wetting the support material in an aqueous solution containing monomer, cross-linking agent and initiator. The membrane-supported gels are then formed through photo-initiation [19] or by heating [20]. The cross-linked polymer chains become entangled within the support matrix, thus stabilizing their morphology with respect to mechanical and osmotic forces [15]. This non-covalent technique results in membrane materials with a high effective concentration of charged groups. An entire chapter of this book has been devoted to the detailed description and applications of gel-filled nanofiltration membranes. 2.2.2
Covalent Attachment Biomolecules
of
Linear
Macromolecules
and
Although cross-linking provides mechanical stability and uniform coverage of the pore cross-section, it also severely restricts chain mobility resulting in poor molecular diffusivity [21]. Sensing devices characterized by
332
Functionalized Membranes For Tunable Separations and Toxic Metal Capture - Bhattacharyya
much faster response times have been developed through surface grafting of linear polymer chains [8]. The immobilization of polymer chains by covalent coupling usually leads to very stable materials with extended life when compared with other coupling methods, namely, physical adsorption and ionic binding. Covalent attachment of linear polymers can be achieved through a variety of functionalization techniques. These methodologies include plasmainduced grafting [8, 22], radiation-induced grafting [13, 23] and chemical grafting [7, 17]. There are significant advantages and drawbacks associated with each technique. 2.2.2.1 Plasma-Induced Polymerization Plasma-induced polymerization is a well-established surface modification technique that can be utilized to form polymer chains within porous supports. The attached polymer is linear because it propagates from a radical formed on the polymer backbone of the support matrix. Radicals formed within the pores of the substrate can be utilized as an initiator for subsequent polymerization [24]. The disadvantage associated with this technique is that plasma species (electrons, ions, etc.) interact strongly with the polymer matrix and with the monomer vapor through various undesirable side reactions [23]. This leads to difficulties in characterization due to the formation of an array of different functionalities. An example of the possible uses of plasma-induced polymerization was the development of a novel, molecular recognition gating membrane using Nisopropylacrylamide (NIPAM, see Figure 1) containing a crown ether receptor [8]. This particular polymer is appropriate for sensing applications because it can recognize a chemical signal, in this case Ba 2÷ ions, and respond through a conformational change exhibiting highly responsive gating properties. Comprehensive reviews on the properties and applications of NIPAM as a stimuli-responsive material can be found in the literature [25-26].
2.2.2.2 Radiation-Induced Polymerization Radiation-induced grafting is a highly versatile functionalization technique. It allows for the attachment of virtually any macromolecule, which can be polymerized through radical addition, with any polymer backbone forming radicals via irradiation (UV radiation, y-radiation or electron beam) [23]. The basis of this method is that ionizing radiation excites radicals in the polymeric backbone of the support material. Upon immersion in monomer containing solution, the polymerization reaction is initiated. The extent of grafting is regulated by the reaction conditions (time, radiation dose, and
333
Functionalized Membranes For Tunable Separations and Toxic Metal Capture - Bhattacharyya
concentration) [23]. In addition, the monomer solution is never exposed to irradiation alleviating problems typically associated with plasma-induced polymerization. One drawback associated with radiation-induced polymerization is that the solubility of the substrate material can be altered by this technique [24]. Stimuli-responsive membranes have been prepared through graft polymerization of acrylic acid (see Figure 1) onto a porous poly(vinylidene fluoride) support via radiation induced grafting [23]. These tunable membranes showed variable permeability upon changes in the pH and ionic strength of the permeate solution. The response or permeability change associated with this particular support was found to be highly dependent on the degree of polymer grafting. The concept of introducing a pH-sensitive response via grafted poly(acrylic acid) (PAA) has been used in the development of novel drug delivery systems [27-28]. In these particular studies, glucose oxidase was coimmobilized through amide bond formation with the grafted carboxyl side chains of PAA. The introduction of glucose into the system resulted in its conversion to gluconic acid resulting in a decrease in local pH. The subsequent protonation of the grafted PAA chains induces a conformational change causing the membrane pores to open and the release of insulin into the permeate. 2.2.2.3
Direct Covalent Attachment
Common problems associated with in-situ polymerization and other functionalization techniques are controlling the degree of polymerization (chain length) and the handling of expensive, poisonous initiators. To overcome these deficiencies, a great deal of research has been directed towards the development of functionalization techniques involving pre-existing, uniform macromolecules. In our work, we have been examining the use of polypeptide functionalized membranes (cellulose, silica) as high capacity sorbent materials [17, 29] and for variable permeability membranes [7]. These membranes were prepared via single-point, covalent bonding of poly(amino acids) to functionalized (aldehyde or epoxide), microporous supports. Synthesis consists of two steps, derivatization of surface functionalities followed by polypeptide attachment under convective flow conditions. The method of derivatization is dependent upon the type of membrane used as the support material. For cellulosic membranes, the formation of surface aldehyde groups was achieved through periodate oxidation [17] or by ozonation [7]. For silica-based membranes, epoxide formation was obtained by silane (3-glycidoxypropyltrimethoxysilane) attachment [ 17]. This reaction involves the methoxy groups of the silane with the silanol groups present on the silica support. Single-point attachment was then performed by permeating aqueous poly(amino acid) solutions through the
334
Functionalized Membranes For Tunable Separations and Toxic Metal Capture - Bhattacharyya
derivatized support under basic conditions. This is done to ensure that the amine terminus of the polypeptide is not protonated and will participate in the nucleophilic attack of the aldehyde/epoxide moities present within the pores of the support. This simple functionalization technique allows for attachment of macromolecules with well-characterized molecular weight distributions. Furthermore, it avoids the use of toxic initiators common to alternative methodologies.
2.2.2.4 Gold-Thiol Chemistry Attachment Techniques The development of comprehensive models describing transport in microporous membranes functionalized with linear polymers has been somewhat limited due to the non-uniform pore distributions common to most support materials (polyethylene, polypropylcne, cellulose, etc.). The fabrication of well-defined stimuli-responsive materials of this type can be achieved through simple gold-thiol chemistry involving track-etched supports [4, 6]. The chemisorption of thiols onto gold-plated surfaces is a well-established phenomena [30]. Deposition of Au nanolayers can be achieved through electrodeposition or by an electroless plating method [4, 30-31]. The latter technique, although more experimentally intensive, has shown more uniform metal deposition. In the electroless method, gold deposition is initiated at the pore walls yielding hollow cylindrical channels at short deposition times [32]. Self-assembled structures of thiol-containing compounds (poly(glutamic acid) with a terminal disulfide group [4], cysteine [6], etc.) spontaneously chemisorb on these gold surfaces. Drawbacks to this method include the formation of "bottleneck" pore structures that would introduce irregularities to the welldefined geometry of the support [30]. Also, practical applications are limited due to the low porosity associated with track-etched membranes and the number of processing steps required in the gold plating procedure. A number of research groups have employed gold-thiol chemistry in the development of environmentally-sensitive membrane materials [4, 6, 30, 32]. One such application has been the chemisorption of L-cysteine onto a goldcoated track-etched polycarbonate membrane having molecular dimensions (pore size as small as 0.9 nm) [6]. The amphoteric properties of chemisorbcd Lcysteine allows for adjustable ion-selectivity of organic anions and cations based on the pH of the feed solution.
335
Functionalized Membranes For Tunable Separations and Toxic Metal Capture - Bhattacharyya
2.3
PROPERTIES OF MEMBRANE MATERIALS CONTAINING TERMINALLY-GRAFTED LINEAR POLY(AMINO ACIDS)
The primary focus of our research group, regarding tunable separations, has been the functionalization of microporous membranes with charged (acidic or basic) poly(amino acid) homopolymers. It has been shown that the morphological and electrostatic properties of these biomolecules can be utilized for selective separations and controlled transport applications [7]. As described in the previous section, attachment involves the reaction of the amine terminus of the polypeptide with functional groups present within the pore structure of the support material.
2.3.1
pH-Sensitive Permeability and Ion-Selectivity
Various polypeptides, such as poly(L-glutamic acid) (PLGA), have been shown to undergo a conformational transition from a random coil structure to a compact helix based on variations in the properties of its contacting solution [4, 7, 33]. These properties include pH, salt concentration and counterion valency [33-35]. In particular, the effects of solution pH on the charge density and morphology of PLGA (see Figure 2) can be used to tune the separation properties of microporous support membranes. The conformation of this macromolecule is highly dependent on the degree of proton binding exhibited by the functional groups present along its backbone. The degree of protonation (X) of the attached polypeptide is given by equation 1 [ 17],
x =
K' exp(-eog/ / kT)[H] +
(1)
1 + K' exp(-eo¢," / kr)[H] + where K' is the intrinsic binding constant (K' = 1/K where K is the intrinsic dissociation constant), eo is the elementary charge, ~g is the electrostatic potential produced by the charged side group constituents, k is the Boltzmann constant, T is the temperature, and [H+] is the hydrogen ion concentration. At pH conditions above 5.5, the degree of proton binding is low and PLGA is fully ionized. Electrostatic repulsive forces between neighboring side groups cause the molecule to unwind and extend into a random coil formation [33-34]. In this conformation, PLGA molecules will extend into the membrane pore in a highly charged state, as shown in Figure 3a. This will allow for electrostatic interactions with ions present in the bulk solution in regions far removed from the pore surface. Also, the presence of these charged macromolecules creates an obstruction to solvent (water) transport. Conversely
336
Functionalized M e m b r a n e s For Tunable Separations and Toxic Metal Capture - Bhattacharyya
at low pH (< 4), a high degree of protonation alleviates the repulsive forces between side group constituents inducing a compact helical morphology as can be seen in Figure 3b. This structure is stabilized primarily by intramolecular hydrogen bond formation. Under these conditions, PLGA offers much less resistance to solvent transport. Furthermore, the reduction of the charge density of the attached macromolecule associated with this conformational shift will result in a dramatic change in the interaction between ionic species and the membrane.
e ChargedSide-GroupConstituent I
a)
|
rp Top View
Side View
b) Core Region Expands i
,,
Figure 3. Schematic representing the helix-coil transitions within the pore of a poly(Lglutamic acid) functionalized membrane; a) random-coil formation at pH > 5.5, b) helix formation at pH (< 4) (rp= pore radius, rc = pore radius - polymer chain length).
2.3.2
Permeation Studies The effects of poly(L-glutamic acid) and poly(L-arginine) functionalization on the pure water transport through a cellulosic-based support membrane are shown in Figure 4. Flux measurements, Jv, for both types of membranes were normalized using the maximum flux, Jv, max. At high pH, the PLGA functionalized membrane displayed a marked decline in the observed water flux. As shown in Figure 3a, PLGA exists in a fully extended random-coil
337
Functionalized Membranes For Tunable Separations and Toxic Metal Capture - Bhattacharyya
formation under these conditions. These extended polymer chains mitigate water transport resulting in the observed decline in permeability [5]. The permeability determined at high pH was 2.31 x 10-4 cm3/cm2 s bar. In its ionized form, this would correspond to greater than an order of magnitude increase in permeability over gel-filled MF membranes found in the literature [3]. At low pH, the flux of the PLGA functionalized membrane increased significantly •
........
...........
•
. . . . . . . .
%°
%
0.8
o
X
I~
E o.6
°o °
"t. . . ........... ~ . . . . . . . . . . . . . . . . . . . . . . . . . . . .
.~
> 0.4 • Poly(L-glutamic acid)
0.2
A Poly(L-arginine)
2
i
i
i
r
i
i
3
4
5
6
7
8
pH
9
Figure 4. The pH dependence of water permeation through a cellulosic membrane functionalized with poly(L-glutamic acid) [7] and poly(L-arginine).
indicating the morphological shift of the attached polypeptide. On the other hand, the membrane functionalized with positively-charged poly(L-arginine) showed virtually no flux response to variations in pH. The guanadino side group of this residue has a pKa of about 12.5. Thus, poly(L-arginine) does not undergo a conformational transition under normal solution conditions. This clearly indicates that the environmentally (pH) sensitive properties of PLGA functionalized membranes is derived from the helix-coil transitions of the attached biomolecule. Results reported in the literature for a gold-coated polyearbonate membrane that had been functionalized with PLGA containing a terminal disulfide group are shown in Figure 5 [4]. The observed flux trend was
338
Functionalized Membranes For Tunable Separations and Toxic Metal Capture - Bhattacharyya
similar for both functionalization techniques indicating that the stimuliresponsive nature of PLGA is independent of the base support membrane. i
i
i
.......... 4, .......... ~.. °. ° % °o
0.8
•o
-,.
X °o
0.8
""'Oo
-)
.e
....~. >
"3
0.4
0.2
1
2
3
4
pH
5
6
7
8
Figure 5. The pH dependence of water flux through a poly(L-glutamic acid) functionalized gold-coated polycarbonate membrane (data taken from [4]).
2.3.3
Charged-based Separations In addition to controlled permeability applications, these membrane platforms can be used for ultra-low pressure charged-based separations. Conventional nanofiltration membranes separate electrolytes based on solute size and charge. NF membranes are generally comprised of monomeric acidic (COO-, SO3) and/or basic (NH3+) functionalities incorporated within a relatively dense (pore diameter < 2 nm) polymer matrix. The charged pore surface establishes an electric potential over its cross-section enabling Donnan-type exclusion of electrolytes. The separation principles of NF can be applied to microporous membranes through extension of the electric potential into regions far removed from the pore surface. This can be achieved through immobilization of charged poly(amino acids) [7]. 2.3.3.1
Feed Concentration Effects on Solute Rejection The dependence of feed concentration on solute rejection, R,
339
Functionalized Membranes For Tunable Separations and Toxic Metal Capture - Bhattacharyya
Cp Cf
R = 1-~
(2)
(where Cp and Cf are the solute concentration in the feed and permeate, respectively) for dilute divalent co-ion solutions (Ca 2+ for poly(lysine), SO42 for PLGA) is shown in Figure 6. The decline in solute rejection at higher electrolyte concentrations is consistent with established transport theory for charged media. It is brought about due to the saturation or shielding effect of counterions on the fixed membrane charge [7]. ~~
~
Positively Charged Membrane 12 solutions)
C 0.8 0
= m
dl=.l
@
1~ "3'
0.6
t¢ 0.4
A-BC (MW 36,400)
m
0
0.2
Negatively Charged Membrane (Na2SO4 solutions) ,
0
i
0.5
,
~
,
1
i
1.5
,
2
mM Divalent Coion Figure 6. Relationship between the rejection of divalent co-ions and feed concentration for cellulosic membranes functionalized with poly(amino acids).
2.3.3.2
Examination of this Non-Steric Mechanism of Ion Exclusion
The separation of ionic solutes with a microporous membrane (0.21 ~tm) can be considered a significant achievement when considering that the hydrated radii of calcium and sulfate ions are 0.412 and 0.379 nm, respectively [36]. Similar experiments performed using neutral 482,000 MW dextran with a molecular diameter greater than 36 nm [37] showed less than 10 % rejection for poly(L-lysine) (75,900 Da) functionalized membranes. Thus, the mechanism of
340
Functionalized Membranes For Tunable Separations and Toxic Metal Capture - Bhattacharyya
separation is obviously non-steric and can be attributed solely to the extension of the membrane charge into the inner portion of pore cross-section. This statement is further supported by the fact that poly(L-lysine) (75,900 MW) showed much greater performance than PLGA (36,400 MW) with respect to solute separation.
A
v e~ 0 in
0.8"
@ 0.6-
I'Y 0.4J je e
~,
0
0.2
e •
,oi,,ee, j,l,l,I I I oleDDI°~=e~lel=jjlq=°°~ee • | 18 B J D BB I J I ~T
I
e
e
i
0
0.2
0.4
0.6
0.8
Chain Length /Ave. Pore Radius Figure 7. Dependence of the solute rejection on the ratio between the poly(amino acid) (PLGA) chain length to the average pore radius for dilute Na2SO4 solutions (0.2 mM) with a permeate flux of 2.5 x 104 cm3/cm2 s (some data taken from [7]).
The chain length, associated with poly(L-lysine) extends much further into the membrane pore structure alleviating the effects of solute leakage in the core region (see Figure 3). To further examine the effects of solute leakage in the core region, rejection studies were performed on PLGA functionalized membranes with varying chain length to pore radius ratios at constant pH (- 6) and flux (2.5 x 10 -4 cm3/cm 2 s). The results of these experiments using dilute (0.2 mM) Na2SO4 solutions are shown in Figure 7. This figure clearly indicates that membrane performance is highly dependent upon the limitation of the core region for this type of functionalized platform [7].
341
Functionalized Membranes For Tunable Separations and Toxic Metal Capture - Bhattacharyya
2.3.3.3
Ion Transport in Poly(amino acid) Functionalized Membranes: Effect of Amino Acid Chain Length Electrolyte transport in charged media is generally described by the extended Nernst-Planck (N-P) equation, Jx,i =
~C i
-Di --~ +Vm Ci
-
O i z i C i F c3(I)
RT
(3)
~x
where Jx,i is the ionic flux in the direction of flow (x), Di is the diffusivity of ion i, Vm is the solution velocity across the membrane, zi is the ion valency, R is the gas constant, Ci is the ion concentration, F is the Faraday constant, and • is the total electrostatic potential within the membrane pore [38-40]. Models based on the N-P equations have been well accepted because they account for the superposition of all three modes of electrolyte transport (namely diffusion, convection and electromigration). It is well established that rejection, R, of electrolytes increases to asymptotic values at higher applied pressures for dense membranes (RO, NF) as is shown in Figure 8. For NF type membranes at low flux, the contribution of diffusion (1 st term of eqn 3) to the overall ionic flux is much more prevalent resulting in lower solute rejection. However, due to its 1.0
__ •
• mm= emmwnmmmnmmllmmmmmmnmmummmmm|
A--.~.4
C 0 0.8 mn "~"~ 0.6
~
R
........ Nanofiltration
0.4
Reverse Osmosis
m
0 01~ 0.2
"--
PP Functionalized MF (Long Chain)
=
PP Functionalized MF (Short Chain)
0.0
Low
AppliedP r e s s u r e
High
Figure 8. Schematiccomparison of polypeptide functionalized microporous membranes with conventional nanofiltration and reverse osmosis.
342
Functionalized Membranes For Tunable Separations and Toxic Metal Capture - Bhattacharyya
narrow pore structure, the electrostatic exclusion (3 rd term of eqn 3) of charged species is maintained by a relatively constant electric potential over the pore cross section. Thus, at high applied pressures, where convection is dominant (2 nd term of eqn 3), sustained electrostatic exclusion is coupled with enhanced water flux resulting in an increase in the observed solute rejection (see Figure 8) to an asymptotic value. In contrast, microporous membranes containing terminally anchored charged poly(amino acids) display these characteristics only over a certain portion of the pore cross section. In regions of the pore structure devoid of charge, the core or inner region (see Figure 3), electrostatic interactions can be assumed negligible. Solute transport in these regions will be unabated by the electric potential established by the grafted polymer. This solute leakage effect will be enhanced at higher volume flux resulting in a decline in solute rejection at higher applied pressures as is indicated in Figure 8. More effective membranes of this type can be synthesized through limitation of this core region. This can be accomplished by using polypeptides with a higher degree of polymerization (chain length, see Figure 8), by enhanced polymer loading or by reducing the pore diameter. The effects of poly(amino acid) chain length on the 1
© E 0.8 O
m e
@ A
0.6
t¢ ¢) 0.4
,4,,,I i
:3 O
P-Lys (75.9 kDa) with CaCI 2 solutions
0.2
/~ PLGA (36.4 kDa) with Na2SO 4 solutions 2.0
i
~
i
i
i
3.0
4.0
5.0
6.0
7.0
P e r m e a t e Flux x 10 4
8.0
(cm31cm 2 s)
Figure 9. The dependence of solute rejection on the permeate flux for polypeptide functionalized microporous cellulosic membranes using dilute (0.15 mM) divalent coion solutions.
343
Functionalized Membranes For Tunable Separations and Toxic Metal Capture - Bhattacharyya
relationship between solute rejection and permeate flux are shown in Figure 9. These experiments were performed using dilute, divalent coion (CaC12 or Na2SO4) solutions of approximately equivalent molar concentration. The cellulosic membrane functionalized with positively-charged poly(L-lysine),PLys (MW = 75,900) showed much greater solute rejection than the shorter poly(L-glutamic acid) (MW = 36,400). In addition, the rejection of CaC12 by the poly(L-lysine) functionalized membrane was much more stable over broad ranges of permeate flux (applied pressure). This can be attributed to the overall size of the attached poly(amino acid) and its ability to occupy larger portions of the pore structure.
2.3.3.4
An Example Application of Poly(amino acid) Functionalized Microporous Membranes for Arsenic Removal A viable application of charged-based separations involving porous supports functionalized with polypeptides is the removal of arsenic containing compounds from contaminated drinking waters. Arsenic is a known carcinogen that is linked to various forms of skin and internal cancers [41 ]. It is commonly 1.0 ¢~ 0.8
0
"~
0.6
<~
21 18
0.4
15
~ ~
~
~
0.2
12
X 6
3
4
5
pH
6
7
0.0 3
4
5
pH
6
7
8
Figure 10. The solute rejection of a single salt arsenate (V) solution (9000 ~tg~ as As) by a silica-based microporous membrane (pore size--- 100 nm) functionalized with poly(Lglutamic acid) (MW - 50 kDa) as a function of feed solution pH. Inset: Effect of pH on the hydraulic permeability (designated as A) shown by the PLGA-Silica membrane used in these solute rejection experiments.
344
Functionalized Membranes For Tunable Separations and Toxic Metal Capture - Bhattacharyya
found in drinking water in two oxidation states, trivalent (As(Ill)) and pentavalent (As(V)). Pentavalent arsenic predominates in surface water due to its enhanced thermodynamic stability [42]. It is typically present in natural waters as a divalent oxyanion (HAsO4 2-) making it an ideal solute for lowpressure ion exclusion. The effect of feed solution pH on the ion exclusion of a 9000 ~g/L (as As) arsenate (As(V)) single salt solution using a silica-based microporous membrane (0.1 ~tm pore size) functionalized with poly(L-glutamic acid) ( M W - 50 kDa) is shown in Figure 10. This would represent a fairly high arsenic concentration with respect to the current EPA maximum contaminant level (MCL) for inorganic arsenic (MCL = 10 pg/L). The observed decline in rejection of the As(V) species with decreasing pH can be attributed to a combination of factors. First, a decline in pH shifts the species equilibrium of As(V) from divalent (HAsO42-) to monovalent (H2AsO4", pKa "~ 6.8) [43]. In addition, the attached macromolecule (poly(L-glutamic acid)) undergoes a conformational transition as the degree of protonation of its side group constituents is enhanced. The effect of this transition on the permeability of the functionalized membrane is given in the inset of Figure 10. At pH 7, the measured solute rejection of As(V) was 93.2 % and the hydraulic permeability (A) of the functionalized membrane was 8.8 x 10-4 cm3/cm2 s bar. Under acidic conditions (pH ~ 4), the hydraulic permeability increased to 19.2 x 10.4 cm3/cm2 s bar. This indicates a change in the secondary structure of the attached biomolecule and a subsequent increase in the effects of solute leakage (see section 2.3.3.3) on membrane separation. This coupled with the shift in species equilibrium (divalent (HAsO42-)=> monolvalent (HzAsO4-)) causes the observed solute rejection at pH 4 to decrease to approximately 29.3 percent.
0
POLYLIGAND FUNCTIONALIZED MEMBRANES FOR METAL CAPTURE
The functionalization of MF membranes with appropriate chelation groups allows for dissolved heavy metal (and anions like nitrate and arsenate) removal from drinking water or high quality water production at low pressure and high throughput rates. It should be noted that the tunable membrane studies presented before dealt with separations based on ion exclusion rather than multivalent (Pb 2+, Fe 3+, etc.) metal capture. For example, to exclude ions such as, HAsO42-, one needs carboxylic functionality (such as, PLGA) where as for capturing these ions requires positively-charged side groups (such as those
345
FunctionalizedMembranesForTunableSeparationsandToxicMetalCapture- Bhattacharyya
associated with poly(lysine) or poly(arginine)) The advantage of MF membranebased adsorbents is that functional groups can be attached to membrane pores
00"
+~
SH
SH
1 Convective flow
Figure 11. Schematicof poly-functionalized membrane sorbents with three different types of functional groups (carboxyl, amine, and thiol).
as polymeric ligands with multiple metal binding sites, rather than as monomeric surface functional groups. Hence, interactions between these groups and the heavy metal ions are very rapid, as sorption proceeds under convective flow and transport resistance is minimized. The overall objective of this research with regards to ion exchange is to create inexpensive sorbents for the entrapment of heavy metal ions at high capacity/rate over conventional ion exchange resins. Our system involves the use of inexpensive, commercially available cellulosic and silica membrane materials that are easily functionalized with polymeric ligands (such as, polyamino acids). A schematic showing various ligands inside a MF type membrane pore is shown in Figure 11. A wide variety of microfiltration membranes can be functionalized with polyamino acids for heavy metal sorption. Cellulosic membranes are generally derivatized with aldehyde groups, and polyamino acids are grafted via a Schiff
346
Functionalized Membranes For Tunable Separations and Toxic Metal Capture
-
Bhattacharyya
base reaction [44]. For a cellulose acetate (CA) base support, the derivatization involved two steps: partial hydrolysis followed by periodate oxidation to create
Cd sorption (0.5 - 1.4 g/g) I I Pb sorption (1.3 - 1.9 g/g) -¢¢ -
4
E
•
0.74
1.14
3.64
6.3
CA Composite CA PE-Silica Pure Cellulose mmoles x 104 Aldehyde / cnf External Membrane Area Figure 12. Metal Sorption Capacities for Cellulosic and Silica Based MF Membrane Sorbents. aldehyde groups. Silica-based membranes have also been functionalized (using silane chemistry) with polyamino acids [17]. Figure 12 shows metal sorption capacities of Cd and Pb (from nitrate salts, 1000 mg/L metals) for polyamino acid (polyglutamic and polyaspartic acids) functionalized cellulosic (Cellulose Acetate, CA Composite and CA, and Pure Cellulose) and silica based (Polyethylene,PE-Silica) MF membranes. Metal sorption (pH 5-5.5, no precipitate formation) is extremely high (>10 meq/g for Pb) compared to conventional ion exchange (1-3 meq/g). This high capacity is due to the availability of a large number of ion exchange groups and the role of counterion condensation [17] in addition to ion exchange. Helix formation (induced by ionic strength, pH, and by sorbed metals) also affects sorption capacity by compressing the chains and further enhancing electrostatic field strength. Anion sorption (nitrate, chromate, etc.) has also been demonstrated with positively charged polyarginine (with amine functionality) functionalized MF membranes, such as 0.9 moles Cr2072 per mole exchange group. The mode of metal capture with carboxyl and amine based poly-amino acids is primarily by ion exchange.
347
Functionalized Membranes For Tunable Separations and Toxic Metal Capture - Bhattacharyya
On the other hand, the use of chelating polyamino acids (such as, polycysteine) containing side-chain thiols (SH groups) provide selective and high capture efficiency for mercury salts [44]. Chelated ion exchangers [45] have wide industrial applications. Various researchers in Japan have also reported the incorporation of IDA (iminodiacetate) functionality and diethylamine groups on membrane supports [46,47]. Denzili et al [48,49] have shown the impregnation of membrane structures with multifunctional dyes for the capture of toxic metals, such as Alkali Blue 6B and Cibacron blue F3GA. Functionalized membrane sorbents offer a wide array of advantages for removing toxic metal cations and anions over conventional ion exchange and other pressure driven membrane processes [ 17, 50-52]. First, MF sorbents offer very low-pressure operation at comparable fluxes for RO/NF membrane separation. Production rates are also competitive with ion exchange, which must contend with pressure drop and channeling across the column. Kinetics for metal sorption is also very high. In competitive sorption technologies like ion exchange and activated alumina, high internal surface areas are provided by microporous (1-10 nm) structures, and hence metals must diffuse to the surface. Convective flow is achieved in MF sorbents, and diffusion mass transfer barriers are virtually eliminated. It should be noted that MF sorbents do have a fixed capacity, and hence will become saturated as any other sorbent and must be regenerated, unlike ion exclusion separation processes. For the regeneration of polyamino (carboxyl) a c i d s - based membrane sorbents, helix-coil transition phenomena also provide an added benefit of selective regeneration of metals because of strong retention of high acidity metals in the helix region [52]. For example, at pH 3 one can selectively recover Cd from a high acidity metal such as, Pb.
4.
ACKNOWLEDGEMENTS
The authors would like to acknowledge NSF-IGERT, and the US EPA, and NIEHS for the financial support of this project. The authors would like to thank Dr. J. Hestekin, Dr. S. Ritchie, Dr. L. Bachas, and Dr. S. K. Sikdar for their highly significant contributions on the metal sorption work. Noah Scherrer (NSF/REU support) conducted some of the experiments involving tunable separations. The authors also acknowledge Osmonics Corporation (for cellulose acetate membranes), and Daramic Corporation (for PE-Silica membranes) for providing some of the base membrane supports used in our work.
348
Functionalized Membranes For Tunable Separations and Toxic Metal Capture - Bhattacharyya
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352
New Insights into Membrane Science and Technology: Polymeric and Biofunctional Membranes D. Bhattacharyya and D.A. Butterfield (Editors) © 2003 Elsevier Science B.V. All rights reserved.
Chapter 17
The design of high performance, gel-filled nanofiltration membranes R. F. Childs* and A. M. Mika
Department of Chemistry, McMaster University, Hamilton, ON, L8S 4M 1, Canada SUMMARY This chapter is concerned with the underlying factors that control the performance of polyelectrolyte gel-filled membranes in nanofiltration (NF) applications. It shows how an understanding of these factors can be used to optimize the performance of this new type of NF membrane. 1.
INTRODUCTION
There is a considerable amount of work currently underway on the development of high performance NF membranes. Typically these membranes are based on a supported, dense, thin-film construct, Fig. l a [ 1]. Major strides have recently been made in understanding the fundamental parameters controlling transport through dense thin-films and this is leading to substantial improvements in membrane performance [2-12].
a)
b)
thin dense separatinglayer-~
Vu,
u u o
ouvJ~U,
t
.......
r
. . . . . . . . . .
Fig. 1. Nanofiltration membrane constructs, a) conventional, b) gel-filled.
353
The Design Of High Performance, Gel-Filled Nanofiltration Membranes - Childs
In our work we have been examining the use of relatively thick, but low density, charged gels anchored within a microporous support membrane as an alternative architecture to thin-films, Fig lb. Membranes with this type of construct can have performances that at least match those of the best NF membranes currently available commercially [ 13,14]. Pore-filled membranes are not new p e r se. For example, there is a considerable amount of work reported on supported liquid membranes [15-18]. These membranes can be effective in a variety of applications; however, they typically lose the incorporated liquid over time and are not stable when subjected to a trans-membrane pressure difference [ 19]. Approaches to overcoming this stability problem include skinning the two surfaces of the membranes thereby physically encapsulating the liquid [20-22]. An alternative and effective approach involves anchoring the separating component within the pores of the microporous host and thus preventing leakage under operating conditions. Typically the separating component is a functionalized, solvent swollen polymer that is either grafted to the pore-walls of the support [23-26] or entangled with the structural elements of the support by cross-linking of the polymeric gels [27-31 ]. In this latter case there need be no chemical bond between the supporting membrane materials and the incorporated gel. Table 1. Water-softening capability of pore-filled membranes obtained with untreated municipal tap water a at 100 kPa. Membrane
Flux
Rejection (%)
kg/m2h
Na +
Mg 2+
Ca 2+
Poly(vinylbenzyl ammonium) (Gel 2)b
12
35
79
80
Poly(4-vinylpyridinium) (Gel 1)b
8
20
74
61
<5
DESAL-51 (Osmonics)
7
21
71
54
>60
Sucrose
a Typical composition: Na+: 12 ppm; K+: 0.8 ppm; Mg2+: 8 ppm; Ca2+: 38 ppm, F: 1 ppm; CI: 25 ppm; NO3: 2.8 p p m ; 8 0 4 2 : 26 ppm; HCO3" and CO32: unmeasured; pH: 7.8. b For structures of the gels see Fig. 3.
The pore-filled membranes used in NF applications contain water-swollen polyelectrolyte gels. The potential performances of this type of membrane are evident when compared with some of the best commercial, water-softening membranes. Comparison of several pore-filled membranes is made with Desal-51, a high performance NF membrane produced by Osmonics, Table 1. As can be seen, pore-filled membranes containing a crosslinked poly(vinylbenzyl ammonium salt)
354
The Design Of High Performance, Gel-Filled Nanofiltration Membranes - Childs
gel have fluxes that are almost twice that of Desal-51 at 100 kPa with somewhat higher bivalent cation rejections. These pore-filled membranes exhibit reasonably linear pressure-flux relationship up to at least 500 kPa and there is no loss of gel or irreversible changes observed at these higher pressures and high flux conditions. Separation changes with flux in a fairly typical manner [14]. The obvious question that arises from these results is what is the NF performance limit of polyelectrolyte filled membranes? In other words, can their performances be substantially enhanced from that shown in Table 1 to the point that they would represent a significant improvement in ultra-low pressure water softening applications? In order to answer such questions it is necessary to fully understand the factors affecting the performance of these pore-filled membranes. 1.1.
Some Basic Considerations
Pore-filled NF membranes consist of two components, namely a support component and a water-swollen polyelectrolyte gel, Fig. 2. The support membrane provides containment and protection of the gel, preventing or restricting its osmotic swelling. The host also largely prevents hydraulic pressure and/or flow induced changes in gel distribution and morphology when in use. Solvent and solute transport through the membrane only occurs through the gel phase.
polyelectrolyte chains
Fig. 2. Schematic representation of the components of a polyelectrolyte-filledmembrane. This simple picture of a polyelectrolyte-gel filled membrane, Fig. 2, already points to an inherent trade off that must be made in membrane design, namely the balance between the porosity of the support membrane and its mechanical strength. In principle, the highest flux would be achieved with no support membrane, i.e.,
355
The Design Of High Performance, Gel-Filled Nanofiltration Membranes - Childs
A100% porosity". This is an untenable situation, however, due to the unrestricted swelling or compression of the gel. On the other hand, a low porosity substrate would afford a membrane with excellent mechanical properties but a low flux as the gel-volume fraction is reduced. A second point to come out of this simple picture is that the nature of the gel itself will be the primary determinant of membrane performance. We can view the gel phase as a two-component system with a gel-polymer and water, Figure 2. Water/solute transport will occur through the water and not the gel-polymer itself. In other words, we would expect the volume ratio of the gel-polymer to water to be a major determinant of flux. Polyelectrolyte gels bear a charge and as such a Donnan potential will be established between the membrane and the surrounding solution. The magnitude of this potential, and hence separation ability of ionic solutes, will depend on the relative charge concentration in the membrane and solute concentration in the contacting solution. Clearly there will be a trade offbetween the requirement for low gel-polymer volume fractions for high flux and higher volume fractions needed to achieve high charge concentrations for good ionic solute rejection. These are very simple considerations of performance. In the following sections we take a more quantitative look at these factors and show how they can be used to obtain polyelectrolyte filled membranes with significantly better performances than those in Table 1. 2.
FACTORS AFFECTING THE PERFORMANCE OF POREFILLED MEMBRANES
2.1
Hydraulic Flux/Permeability Determinants: the Properties of the Gel As noted above, transport of solvent through gel-filled membranes occurs only through the gel phase. Kaput and Anderson [27] have shown that the Darcy permeability ofpoly(acrylarnide)-filled membranes is related to the permeability of the pore-filing gel, as shown in Eq. (1) k m =kp £
(1)
where km is membrane permeability, kp is gel permeability, e is the porosity of the substrate and r, tortuosity. The tortuosities of typical microporous membranes do not differ greatly, however, their porosities are an important determinant of
356
The Design Of High Performance, Gel-Filled Nanofiltration Membranes - Childs
membrane permeability. As noted above, the substrate should have as high porosity as is practicable and consistent with membrane strength. The intrinsic permeability of the gel is the dominant control parameter. There are a large number of studies on gel permeability including gel sedimentation rates. In essence, gel permeability is a function of the gel-polymer volume fraction, ~. The same relationship holds when the gels are constrained from swelling by incorporation into a microporous substrate, however, one must be careful to base the gel-polymer volume fraction on the pore volume of the support membrane and not that of an unconstrained gel.
kp = f ( ~ )
(2)
We have studied the relationship between membrane permeability (or gelpermeability through Eq. (1)) and the polymer volume fraction of a variety of polyelectrolyte gels incorporated into the same poly(propylene) host membrane.
"
1
2
HO~~,.OH
3
H HO= ~O,H''~" "~o.. 4
5
..~-o
me~O.,H E
Fig 3. Structures of polyelectrolyte gels inserted into microporous hosts: 1 - poly(4vinylpyridinium salt), 2 - poly(vinylbenzylammonium salt), 3 - poly(ethyleneimine), 4 poly(acrylic acid), 5 - poly(styrenesulfonic acid), 6 - poly(2-acrylamido-2methylpropanesulfonic acid).
As shown in Fig. 3, these gels include positively charged gels (anion exchange membranes) and negatively charged gels (cation exchange membranes). The relationship between gel permeability and gel polymer volume fraction for membranes containing polymer 1; (R = N-benzyl) is shown in Fig. 4. Over the polymer volume fraction studied, there is an inverse relationship between ~ and kp 357
The Design Of High Performance, Gel-Filled Nanofiltration Membranes - Childs
that follows the relationship shown in Eq. (3). kp = 6.10 x 10- 2 2 ~ - 3 . 5 5 (m z)
(3)
There is a very strong dependence of permeability on polymer volume fraction for this gel polymer and it is only at low values of ~bthat membranes will have high permeabilities. This relationship shows why typical ion-exchange membranes, which have low water contents and consequently high polymer volume fractions, are not suited for pressure driven applications.
100
A
10
om
.Q t~
E
'-Q I1.
1
"~ (.9
Calculated
0.1
.....
0~
011
Gel-Polymer Volume Fraction ~
(-)
Fig. 4. Relationship between gel Darcy permeability and gel-polymervolume fraction for a membrane made using gel 1. The solid line is derived using the model described below. (Reproduced with permission from Ind. Eng. Chem. Res. 40 (2001) 1694-1705. Copyright 2001 Am. Chem. Soc.). We have recently described a model that can be used to calculate gel permeability from first principles [32]. The gel is modeled as an assemblage of spheres, Fig. 5, with the sphere diameter being equal to the correlation length, (, of an equivalent semidilute polymer solution. The correlation length is calculated according to the Schaeffer relationship [33], Eq. (4), where n is the number of
358
The Design Of High Performance, Gel-Filled Nanofiltration Membranes - Childs
bonds of the length a in the polymer persistence length, X is the Flory-Huggins interaction parameter, and w is the three-body excluded volume parameter.
n2a ~=[n(l_2z~+wa_6¢2In
(4)
The effect of charge on the persistence length of the polyelectrolyte is included. The permeability of the sphere assemblage is calculated using Happel=s cell model [34].
Fig. 5. Gel modeled as an assemblage of spheres. It can be seen from the solid line shown in Fig. 4, this sphere model for gel permeability predicts the experimental data exceptionally well. Other models for gel permeability, such as the fiber model [35,36], cannot be used to fit data obtained over a wide range of polymer volume fractions without the artificial device of adjusting the thickness of the fiber. We have examined the importance of the various terms in the Asphere@ model. The gel-polymer concentration is a dominant factor. In principle chain stiffness, n, is important and it would be expected that this would be greatly affected by the high nominal charge on the polyelectrolyte-gels. However, this effect is not of major importance due to charge shielding and counter-ion condensation. The Flory-Huggins solvent interaction parameter is also an important term. This is exemplified in Fig. 6 where the permeabilities of two gels with 359
The Design Of High Performance, Gel-Filled Nanofiltration Membranes - Childs
different solvent interaction parameters are calculated as a function of ~b. It is interesting that the more hydrophilic the gel polymer, the lower the permeability of the gel for a given value of ~b. A
E
100
"7,
•~-
lO
"'''"
Q
0.45
E
a.Q
1
e-
.m
~0.3s_J ~
.....
,10
E Q ~
0.1
O.O4
0.08
0.12
0.16
Gel-Polymer Volume Fraction, ~ (-)
Fig. 6. Effect of the Flory Huggins polymer-solventinteraction parameter of the gel on membrane permeability. These effects were verified by measurement of the permeabilities of a series of membranes containing the gel polymers 1-4 and 6, Fig. 7. A series of correlations is seen for these membranes with the permeabilities being dependent on the gel polymer. The calculated permeabilities of the membranes, shown as the lines in Fig. 7, closely match the experimental results. Poly(vinylpyridine) and poly(vinylbenzyl ammonium) based systems ( gels I and 2, respectively) have very similar permeability relationships. Poly(acrylic acid) (gel 4) filled membranes have permeabilities slightly lower than those with gels 1 and 2, consistent with higher hydrophilicity of 4. Poly(ethyleneimine)-based membranes (gel 3) show somewhat higher permeability caused by the higher rigidity of the branched polymer. The results obtained with poly(2-acrylamido-2-methyl-propanesulfonic acid) gels, 6, show the largest difference from the other gel-polymers and membranes made with this gel-polymer have significantly lower permeabilities. This is in large part due to the very hydrophilic nature of this polymer. At this point we cannot model this polymer precisely as we are in process of determining some of its basic parameters to insert into the model. It can be seen, however, that the shape of the relationship between membrane permeability and polymer-volume fraction parallels those of the other systems. The assumption made in the Asphere@ model presented above is that the 360
The Design Of High Performance, Gel-Filled Nanofiltration Membranes - Childs
incorporated gels are homogeneous and evenly distributed through the thickness of the membrane. For all the membranes shown in Fig. 7, we carefully checked that when the gels were made in test-tubes they were perfectly clear and showed no solvent syneresis. If this is not the case then deviations from these correlations would be expected to occur. This can be seen in Fig. 8 where the permeabilities of a series of membranes with a constant polymer fraction of a poly(vinypyridinium salt) gel but increasing degree of cross-linking are shown. ,,,10000
E ~" "7. O
~e
1000
>~
100
m ... "Q t~
10
:" ",~: ~,
4)
'~'~,.'.:,., .
E e~
[] [] ....... [] .......... • ©
D[]
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1- Model 1- Experiment 2 - Experiment 3 - Model 3 - Experiment 4 - Model 4 - Experiment 6 - Experiment
I
o.1
E O.Ol 0.001 0.00
0.02
0.04
0.06
0.08
0.10
0.12
0.14
0.16
0.18
0.20
Gel-Polymer Volume Fraction, # (-) Fig 7. Membrane permeability as a function of ~bfor various gel polymers (for chemical structures of gel polymers see Fig. 3).
As can be seen, counter to one's intuition where one would expect a reduction in permeability as the net size of the gel decreases with increasing crosslinking, there is a sizeable increase in permeability. The origin of this effect would appear to lie in increasing heterogeneity being induced in the gel with increasing cross-linking. This will have the effect of forming polymer rich ~on-draining@ and polymer depleted Adraining@ regions within the gel. Given the large sensitivity of permeability to the polymer volume fraction, the effective reduction in pore volume due to the formation of the non-draining regions is more than compensated by the reduction in polymer volume fraction of the draining regions leading to the enhanced permeability. Similar effects are seen with poly(acrylic acid) gels where micro-heterogeneity in the gels can be induced by using a poorer solvent in the polymerization reactions used to make the membranes [37]. It is interesting to note these effects can also be dynamic. We have recently shown that reversible changes in permeability of membranes can be induced by the use of weak base polyelectrolyte gels such as poly(4-vinylpyridine) and feed solutions of different 361
The Design Of High Performance, Gel-Filled Nanofiltration Membranes - Childs
pH [38]. The overall picture emerging from this work is that the permeability of these gel-filled membranes is very predictable if the properties of the gel are known. More work is required to fully define all the parameters, particularly the quantitative analysis of gel heterogeneity and the effect of counterion, and water of solvation on permeability. However, this analysis already provides the quantitative tools to optimize the properties of the gel. Clearly these ultimately must be combined with the separation properties associated with these gels and this will be analyzed later in this chapter. But first, how does the base membrane affect these membranes? A N
100
Degree
of Crosslinking
(mol-%):
0 t25 0,-20 e,..15
G. r,,
..5-10
1
.Q
E o :E
0.1 0.04
i ,
,.
, i 0.06
,
,
,
,
i
0.08
,
,
,
J i 0.1
,
,
,
J i 0.12
,
,
,
,
i
0.14
. . . . 0.16
Gel-Polymer Volume Fraction, ~ (-) Fig 8. Effect of increasing the degree of cross-linking on the permeability of a membrane containing a poly(4-vinylpyridinium salt), gel 1 (line: model, points: experiment).
2.2 Hydraulic Flux~ermeability Determinants: the Properties of the Supporting Membrane The above analysis of membrane permeability has largely focused on the properties of the incorporated gel and the importance of the base membrane parameters such as porosity and pore tortuosity. As hydraulic flux under pressure is a key operating parameter, it is important to know what other factors are important in selection of a suitable supporting membrane. A comparative study of two types of support membranes has been carried out. The first support is a poly(propylene) membrane produced by the 3M Company using a thermally induced phase separation process (TIPS), Fig. 9a, and the second one is a non-woven, ultra-high molecular weight poly(ethylene) membrane produced by DSM Solupor, Figure 9b. Examination of the SEM images 362
The Design Of High Performance, Gel-Filled Nanofdtration Membranes - Childs
in Figure 9 shows that these membranes have quite different morphologies. Despite this, their porosities are essentially identical (80%). The thicknesses of the starting membranes are given in Table 2. It should be noted that the 3M and first Solupor membranes have essentially the same thickness. The second Solupor membrane has a much smaller thickness.
•
Fig 9. SEM micrographs of different microporous support membranes: a) 3M Company poly(propylene) and b) DSM Solupor UHMW poly(ethylene) membranes.
The pure-water flux of a pore filled membrane, Jw, is related to the membrane permeability, kin, thickness, Ax, of the gel layer, and the pressure difference, AP, across the membrane, as shown in Eq. (5). AP J w oE k m ~
(5)
Ax
The gel layer occupies the entire thickness of the support membrane in the porefilled membranes currently under discussion and, as a result, the gel thickness corresponds to the total average thickness of the pore-filled membrane. Thus, if one continues to fill the entire thickness of the membrane with the gel then to enhance flux we need to use thinner, but still high porosity, support membranes. However, the issue is more complex than this as is shown by the results obtained with the two 363
The Design Of High Performance, Gel-Filled Nanofiltration Membranes - Childs
hosts. The two supporting membranes were filled with either poly(N-benzyl-4vinylpyridinium salt) (1 R=Be) or poly(N,N,N-trimethyl-vinylbenzylammonium salt) (2) gels with very similar gel-polymer volume fractions. The thicknesses of the resulting gel-filled membranes were measured as well as their pure-water fluxes. The results are given in Table 2. Table 2. Comparisonof different host-membrane substrates. Basemembrane
Filling gela
3M
Thickness (,um)
Flux at 100 kPa (kg/m2h)
Starting
Filled
Measured
Calculated
2
79
80
7.3
7.1
Solupor
2
92
146
3.8
21.4
Solupor
2
47
66
8.2
13.7
Solupor 1 (R=Be) 47 aFor structures of the gels see Fig. 3.
160
5.2
18.5
Several key features stand out from these results. First, there is a substantial difference in the ability of the two different types of membrane to resist osmotic swelling of the gels. This is clearly evident from the first two entries where the nonwoven Solupor membrane undergoes a substantial increase in thickness on introduction of gel 2 while the 3M support remains the same thickness within error. The degree of thickness increase also depends on the properties of the gel-polymer chosen. This is evident in the last two membranes where the more hydrophilic poly(vinylpyridine) causes the base membrane to swell by more than a factor of 3! The pure-water fluxes (100 kPa) of the membranes were measured and compared to the values calculated using the permeability model referred to above, Table 2. As the exact masses of incorporated gels are known the polymer volume fractions, membrane permeabilities, and membrane fluxes can easily be calculated using the measured dimensions of the various gel-filled membranes. As would be expected based on the substantial thickness increase associated with the Solupor membranes, they are predicted to have substantially lower gel-polymer volume fractions and, consequently, higher fluxes than the 3M support-based membrane. The agreement between the measured and estimated fluxes of the 3M-based membrane is remarkable and points to the power of the permeability model outlined above. The results with the Solupor membranes are quite different, and in each case
364
The Design Of High Performance, Gel-Filled Nanofiltration Membranes - Childs
the measured fluxes are substantially smaller than the estimated values. It would appear that just as the Solupor membranes lack the necessary mechanical strength to prevent osmotic swelling of the membrane, they also lack the strength to prevent compression of the membrane under hydraulic flow. Such a compression would increase the effective gel-polymer volume fraction and reduce flux. Evidence supporting such a compression comes from the pressure/flux relationships for the Solupor based membranes where, unlike the ones typically found for the 3M membranes, they are markedly non-linear [14]. It should be stressed that as membrane permeability is so strongly dependent on the polymer volume fraction of the gel, even small compressions of the membrane will lead to very significant reductions in flux. One further feature of the membranes summarized in Table 2 is the middle two entries where the membranes have identical gel compositions and concentrations but the support membranes have different thicknesses. Just as would be expected, the thinner membrane has a substantially higher flux. Overall, these results with the two different supports emphasize the importance of choosing a support with the appropriate properties. The Solupor membrane is a very interesting product, however, it is not well suited for use with polyelectrolyte gels in that it lacks strength in the thickness dimension. The 3M TIPS membrane couples a high porosity with high mechanical strength. We also note that care must be taken not to modify the properties of the base membrane during insertion of the gel. It would appear, for example, that with membranes made by in-situ radical polymerization of monomers there is frequently some modification of the support presumably resulting from either grafting or interpenetration of the gel polymer with the support. This can lead to fairly large thickness changes in the membranes on introduction of the gel. On the other hand, membranes made by in-situ cross-linking of a preformed polymer typically show greater dimensional stability. 2.3
Determinants of Separation: Neutral Solutes The results given in Table l show that there is a key difference in separation performance between polyelectrolyte gel-filled and conventional thin-film membranes. This difference lies in their basis of separation. For example, with sucrose, a relatively large organic molecule, a Desal-51 thin-film membrane exhibits a fairly large rejection (>60% at 100 kPa) while a poly(4vinylpyridinium salt) gel based membrane has <5% rejection at the same pressure. While thin-film membranes such as Desal-51 normally have a negative
365
The Design Of High Performance, Gel-Filled Nanofiltration Membranes - Childs
surface charge, their separation mechanism includes a sieving term and hence the rejection of neutral molecules such as sucrose. In gel-filled membranes, on the other hand, the Apore@ diameter is equal to the mesh size of the gel. In terms of the Asphere@ model of gels described above, the gel mesh size is formally equal to the correlation length, ~, and thus dependent on the gel polymer volume fraction, ¢L This is illustrated by the simulation of the asymptotic rejection of polyethylene glycols (PEGs) of different molecular weights as a function of gel-polymer volume fraction, Fig. 10. The rejection curves were calculated from the steric hindrance model described by Bowen et al. [9] using 2~ as the pore radius, rp. The StokesEinstein radii of PEGs and the bulk diffusion coefficients were calculated from the intrinsic viscosity of the polymers, according to Singh et al. [39]. Gel-filled membranes with fluxes comparable to those of commercial thin-film composite membranes typically have gel-polymer volume fractions in the range of 0.08#¢~ #0.12 and, on the basis of the data shown in Figure 10, such membranes should show a >90% asymptotic rejection of PEG 4000.
1.0 .s"
0.9
i Em~O. 8
•"
0.7
s "S
..o.....
~'"
0 0.6
J •
.....
"""•"
/"
.o"*"
•.......
m 0.5
O:EE!:OOO
"~'rw 0.412
"""'°""
0.3
~.
0.2
E
o.1
]
o.o . . . . 0.02 0.04
0.06
0.08
0.10
0.12
0.14
0.16
0.18
0.20
Gel Polymer Volume Fraction, ¢ (-)
Fig. 10. Asymptotic rejection of PEGs of different molecular weight as a function of poly(4vinylpyridinium salt) volume fraction.
366
The Design Of High Performance, Gel-Filled Nanofiltration Membranes - Childs
1.0
'm 0.8
0.6
'0.4
n,
Legend
~ o . 2 tl
- -
•
Mode!
Experiment
0.0 0.0
1.0
2.0
Volume
3.0
4.0
5.0
6.0
Flux, Jv (104 m/s)
Fig 11. Calculated and measured rejection of PEG 4000 by a poly(N-methyl-4-vinylpyridinium chloride) membrane with ~b= 0.08V0.005.
This prediction of PEG 4000 rejection, based on the assumption that rp .2~, has been tested experimentally. The rejection of PEG 4000 obtained with a membrane with a poly(N-methyl-4-vinylpyridinium chloride) gel (polymer volume fraction ~b= 0.08V0.005) is shown in Fig. 11 as a function of volume flux. The real rejection of a PEG 4000 was calculated from the observed rejection using the mass transfer relations for a stirred cell [9]. As can be seen the real rejections reach a value of ca. 0.8. The steric hindrance model [9] was used to calculate the fluxrejection curve for a membrane having the thickness and porosity equal to those of the gel-filled membrane. As can be seen from Fig. 11, the experimental points fall very close to the simulated curve. This result further substantiates the modeling approach we have taken with these membranes.
2.4
Determinants of Separation: Charged Solutes The basis of salt separation with these charged, gel-filled membranes is Donnan exclusion rather than sieving. This is well exemplified by the data shown in Fig. 12 where the single salt rejections of NaC1 and Na2SO4 with a positively charged membrane based on a poly(vinylbenzyl ammonium salt) gel are plotted against solute concentration. Both salts have the same monovalent cation (Na+), the co-ion in this case, but different counter-ions. The Donnan potential with the divalent counter-ion (sulfate) is expected to be substantially lower than that with monovalent counter-ion (chloride) and decrease rapidly with increase in solute concentration. As a result, the rejection of sodium sulfate is markedly lower than that of sodium chloride and rapidly decreases with increase in concentration. The question naturally arises as to whether there are suitable models to account for such 367
The DesignOf High Performance, Gel-FilledNanofiltration Membranes- Childs
behavior and which can be used to optimize gel composition? ........................................................................... _ ~
100
90 so
•"
CI
30 a2SO4
20 10 0 0
1O0
200
300
400
500
600
700
800
Salt Concentration, c, (ppm)
Fig 12. Single salt rejection as a function of salt concentration in feed for a membrane with a poly(vinylbenzyl ammonium salt) gel (gel-polymer volume fraction ~ = 0.16; applied pressure: 500 kPa).
The fixed-charge or Teorell-Meyer-Sievers (TMS) model [40-43] predicts that the salt rejection by charged membranes depends on the ratio of the membrane effective charge density to the salt concentration in the feed and not on their absolute values [44,45]. The main assumption underlying the model is homogeneous distribution of aqueous interstices through which convection and diffusion take place. The application of the TMS model to gel-filled membranes is justifiable provided that the distribution of electrical potential and ion concentration across the gels are uniform. Clearly, for gels with large mesh sizes this assumption will be invalid as the concentration of ions will vary in the radial direction. What constitutes a large mesh or pore size? In fact, a radial variation would be expected when the pore radius is substantially larger than the Debye-Htickel screening length, x-:, at the ionic strength of the feed solutions [44]. For such cases, the socalled space charge (SC) model of electrolyte transport through charged capillaries proposed by Osterle and coworkers [46-48] would be a more realistic model to predict salt rejection by porous charged membranes. However, for the NF membranes, under discussion in this chapter, the TMS model is applicable. Thus uniform ion-concentrations and electrical potentials would be expected for membranes with gel-polymer volume fractions in the range of 0.08 to 0.12, provided the ionic strengths of the feeds do not exceed 15 to 30 mM, respectively. The total charge density in a unit membrane volume can be calculated based on the polymer volume fraction in the pore-filling gel, partial molar volume of the gel polyelectrolyte, and porosity of the substrate. The effective membrane charge density may, however, be substantially lower than the nominal densities calculated
368
The Design Of High Performance, Gel-Filled Nanofiltration Membranes - Childs
from the concentrations of ionic sites on the polymer chains due to binding (association) of a fraction of counter-ions to the fixed charge of the polyion. Assuming that no specific binding of counter-ions to the fixed charge takes place, the effective charge density can be estimated from the Manning counter-ion condensation theory [49]. The theory postulates the existence of a limiting value of the charge density along a polyelectrolyte chain above which the system (polyion and counter-ions) is unstable. At this point a sufficient number of counter-ions must condense onto the polyion to reduce its charge density to the limiting value such that the distance between charges along the chain contour is equal to the Bjerrum length (0.713 nm in water at 25°C). For polyelectrolyte chains of vinyl monomers in water, this condensation should start at approximately 35 % ionization [50]. The asymptotic rejections of NaC1 of various concentrations were calculated as a function of the gel-polymer volume fraction using the TMS model [44,45]. The results obtained with positively charged gels ofpoly(4-vinyl-N-benzylpyridinium chloride) are shown in Fig. 13. The dotted lines indicate the regions of gel-polymer volume fractions where the assumptions of the TMS model may be invalid due to the pore-radii becoming larger than the Debye-Hiickel screening lengths. The data presented in this figure show clearly that monovalent salts rejections well exceeding 50 % are available even at the feed concentration as high as 25 mM ( -~ 1500 ppm NaC1) with the higher gel-polymer volume fractions. Rejection is a function of the charge density of the polyelectrolyte gel incorporated into the membranes. Increasing the charge density by increasing the gel-polymer volume fraction, as was done in the data recorded in Fig. 13, will inevitably lead to a reduction in flux at a constant pressure. An alternative way of increasing the charge density would be to use polyelectrolytes with lower molecular weight repeating units and low partial molar volumes. Such an approach should give higher charge densities without a major increase in gel-polymer volume fraction. The net result should be higher rejections without undue reduction in the membrane permeability or flux. We have examined the separation of sodium chloride with the gel-filled membranes containing different pore-filling gels and have found a remarkable similarity [51 ]. In Fig. 14 we show NaC1 rejections for a series of membranes as a function of membrane hydrodynamic permeability. Each point in the figure represents a different membrane. The line is a polynomial fit of the experimental points. The striking feature is that all the membranes, whether filled with strong base or strong acid polyelectrolyte-gels, show directly comparable performances. Indeed, the data can be correlated with a single line!
369
The Design Of High Performance, Gel-Filled Nanofiltration Membranes - Childs
1.00
0.95
0.90 0.85 0.80
0
0.70 0.65
NaCl
0.55 0.50
,
,
,
0.04
,
i
,
,
,
0.06
,
L
,
,
,
0.08
,
i
. . . .
0.1
ncentration
ii i
,
,
,
0.12
mM
,
i
,
,
,
,
0.14
0.16
Polymer Volume Fraction, ~ (-)
Fig. 13. Asymptotic rejection of NaC1 as a function of gel-polymer volume fraction in poly(N-benzyl-4-vinylpyridinium chloride). The dotted parts of the curves are regions where the TMS model may be invalid as described in the text.
100
o~"
©
80
~
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, , ,,~v-,~G~
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Membrane Permeability, k m (10 "18 m 2)
Fig. 14. NaC1 separation for a series of membranes as a function of membrane permeability. (The data was obtained with 300 ppm (-~ 5 mM) NaC1 feeds at 300 kPa). The data in Fig. 14 can be divided into two distinct regions with the boundary being at a permeability of about 2 -_-1018 m 2. This boundary point corresponds to a gel concentration of ~b- 0.08. In the region of lower permeability (~ 3 0.08), the salt rejection is very high (70-80%) but remarkably, this rejection
370
The Design Of High Performance, Gel-Filled Nanofiltration Membranes - Childs
shows tittle sensitivity either to large changes in membrane permeability (over one order of magnitude) or the accompanying changes in charge density. This lack of sensitivity comes from high charge density as compared to the feed concentration. As shown in Figure 13, at 5 mM NaC1 concentration the asymptotic rejection shows little dependence on polymer volume fraction. This means that the ratio of the fixed charge concentration to the salt concentration is high over the whole range of gel concentration. In contrast, rejections in the region of higher permeability (~< 0.08), decrease very rapidly with increasing membrane permeability and approach zero at the permeability of about 10-17m 2. More work is required to fully characterize the behavior of these membranes; for example, the role of factors such as concentration polarization with membranes of very high fluxes and hydraulic instability of gels at low polymer volume fractions need to be probed. However, even at this point, it is clear that the membranes shown in Fig. 14 can readily be optimized in terms of flux and separation by adjustment of the gel-polymer volume fractions.
lOO
90 80
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60
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30
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ly(acrylicacid)
~er
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,
,
i
,
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~
,
,
LL,I
1
Membrane Permeability,
~
,
,
,
10 k m (10 -18 m 2)
a
~
i,i
100
Fig. 15. NaC1 rejection as a function of membrane permeability of poly(acrylic acid) containing membranes (feed: 300 ppm NaC1, pressure: 300 kPa).
The separation behavior of membranes containing poly(acrylic acid) gels is quite different from the general trend found with both strong base and strong acid membranes, Fig. 15. The salt rejection of this weak polyacid is substantially lower than that found with other gels and is independent of membrane permeability (gelpolymer volume fraction) over an extended range. While we continue to study this and related systems, the lower rejection of NaC1 indicates that the effective charge
371
The Design Of High Performance, Gel-Filled Nanofiltration Membranes - Childs
density in these membranes is substantially lower than in the other membranes. This is suggestive of a very strong polyelectrolyte effect and a large shift of its pKa to make it a weaker acid on incorporation into a microporous support. A large change of dissociation constant, amounting to more than two units ofpKa, has been found, for example, for membranes filled with poly(4-vinylpyridine) gels [38,52]. The work presented thus far has largely focused on single salt separations and understanding the behavior ofpolyelectrolyte filled membranes. The situation becomes more complex with mixtures of solutes particularly where specific interactions of solute ions with a gel can occur. For example, the performance of a series of different membranes with a municipal tap-water feed is given in Table 3. As would be expected with separation being based on Donnan exclusion the charge type of the filling gel is important. As can be seen, the negatively charged gels show substantially reduced rejections when compared to positively charged membranes. The lower rejection by the polyacid containing membranes is not only caused by the presence of divalent counter-ions Ca 2+ and Mg 2+ but also by strong specific binding of cations to carboxylic and sulfonic acid groups. This binding has an effect of reducing the effective charge density. The divalent anions SO42- and CO32, the counter-ions for the positively charged membranes, are also present in this feed but their binding to cationic polyelectrolytes is weak and does not significantly affect rejection. Further work is underway to probe these effects more deeply. Table 3. Softening of tap water a (test conditions: untreated municipal tap-water as a feed; stirred dead-end cells; pressure: 100 kPa). Membraneb
Charge
Flux
Rejection (%)
kg/m2h
Na +
Mg 2+
Ca 2+
Poly(vinylbenzyl ammonium) (Gel 2)
+
12
35
79
80
Poly(4-vinylpyridinium) (Gel 1)
+
8
20
74
61
Poly(AMPS) (Gel 6)
-
12
12
22
27
Poly(acrylic acid) (Gel 4)
-
5
10
29
38
7
21
71
54
DESAL-51 (Osmonics)
a For tap water composition see footnote at Table 1; b For structures of the gels see Fig. 3.
3.
CONCLUSIONS In this chapter we have described a new family of NF membranes consisting 372
The Design Of High Performance, Gel-Filled Nanofiltration Membranes - Childs
of a polyelectrolyte gel anchored within the pores ofa microporous host membrane. At the outset of the chapter we pointed out that these Aion-exchange@ membranes have NF performances which match commercial NF membranes and asked the question as to whether their performance could be enhanced. In answering this question attention has been particularly focused on understanding the principles governing the performance of gel-filled membranes. It has been shown that the permeability and flux of these membranes can be calculated with excellent precision using an a-priori model for gel permeability that we have developed. The ability of this model to predict permeability and flux is such that experimentally observed deviations with certain membranes indicate the onset of factors such as gel microheterogeneity or pressure induced compression of the membranes in NF applications. The gel-mesh radii derived from the model can also be used to predict the separation of large neutral solutes. It has also been shown that the separation of ionic solutes is based on Donnan exclusion. Salt rejections can be calculated using the TMS model and a good fit between observed and calculated values are found. Table 4. Some partially optimized NF membranes (feed: municipal tap-water; stirred deadend cells). Membrane
Pressure (kPa)
Flux (kg/m2h)
Na÷
Rejection (%) Mg2+
I
50
18
17
63
61
II
100
35
16
63
61
III
100
38
18
61
64
7
21
71
54
DESAL-51 (Osmonics) 100 a For tap water composition see footnote at Table 1
Based on the principles presented here work is currently underway to fabricate higher-performance NF membranes that will function effectively at ultralow driving pressures. This work is bearing fruit and it is clear that substantial improvements in performance can be achieved. This is evident from the data given in Table 4 and comparing these data with those given in Table 1. The striking features of the various membranes listed in Table 4 are that not only can substantial improvements in performance be made but that it is possible to fabricate membranes capable of softening municipal tap-water with a trans-membrane pressure of only 50 kPa. The polyelectrolyte-filled membranes described in this chapter are not just
373
The Design Of High Performance, Gel-Filled Nanofiltration Membranes - Childs
laboratory curiosities. They are robust membranes despite the intuitive impression given by a membrane construct in which a relatively soft, water-swollen gel is inserted into a microporous support. For example, several membranes have been continuously tested for well over a year in our laboratories with no loss in properties. They are easy to clean, and potentially both simple and cheap to fabricate on a continuous basis. They even come with a Abuilt-in@ self-repairing capability associated with the soft-gel separating component. The ability of the membranes to achieve comparable fluxes and separations but operate at lower pressures offers the prospect of significant reduction in membrane fouling and lower energy consumptions. This work clearly demonstrates that consideration be given to gel-filled membranes as an alternative to the conventional thin-film approaches to high performance membranes. ACKNOWLEDGMENTS The work described in this chapter has been ably carried out by a variety of people at McMaster University. These include Drs A. K. Pandey, H. Byun, graduate students W Jiang, C. McCrory, J. Zhou, and M. Kim, and technical support ofE. H. M. Lee. Financial support from the 3M Canada Corporation, the Natural Science and Engineering Research Council of Canada is gratefully acknowledged. Membrane supports were donated by 3M Canada Corporation and DSM Solupor.
REFERENCES
[1] [2] [3] [4] [5] [6] [7] [8] [9]
[10] [11] [12] [13] [14]
R.W. Baker, Membrane Technology and Applications, McGraw-Hill, New York, 2000. R.F. Childs, Membrane Quarterly, 16(2) (2001) 7. A.M. Mika and R.F. Childs, In: Membrane Separations, A. Noworyta and A. TrusekHolownia (Eds.), Agencja Wydawnicza "Argi", Wroclaw, 2001, p.23. S.-Y. Kwak, C.K. Kim, and J.J. KJm, J.Polym.Sci.:B.Polym.Phys. 34 (1996) 2201. S.-Y. Kwak, Polymer 40 (1999) 6361. S.-Y. Kwak and D.W. Ihm, J.Membr. Sci. 158 (1999) 143. S.-Y. Kwak, S.G. Jung, Y.S. Yoon, and D.W. Ihm, J.Polym.Sci.:B.Polym.Phys. 37 (1999) 1429. W.R. Bowen and H. Mukhtar, J.Membr. Sci. 112 (1996) 263. W.R. Bowen, A.W. Mohammad, and N. Hilal, J.Membr. Sci. 126 (1997) 91. W.R. Bowen and T.A. Doneva, J.Colloid Interface Sci. 229 (2000) 544. W.R. Bowen and T.A. Doneva, Desalination 129 (2000) 163. P. Fievet, A. Szymczyk, B. Aoubiza, and J. Pagetti, J.Membr. Sci. 168 (2000) 87. A.M. Mika, R.F. Childs, and J.M. Dickson, Desalination 121 (1999) 149. A.M. Mika, A.K. Pandey, and R.F. Childs, Desalination, 140 (2001) 265.
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[15] [16] [17]
[18] [19] [20] [21] [22] [23] [24] [25] [26] [27] [28] [29]
[3o] [31] [32] [33] [34] [35] [36] [37] [38] [39]
[40] [41] [42] [431 [44] [45] [46] [47]
[48] [49]
[50]
G. Schulz, Desalination 68 (1988) 191. F. Nakashio, M. Goto, and T. Kakoi, Solvent Extr.Res.Dev.,Jpn. 1 (1994) 53. F. de Jong and H.C. Visser, Comp.Supramol.Chem. 10 (1996) 13. M. de Gyves and E.R. de Miguel, Ind.Eng.Chem.Res. 38 (1999) 2182. A.J.B. Kemperman, D. Bargeman, Th. van den Boomgaard, and H. Strathmann, Sep.Sci.Technol. 31 (1996) 2733. M.C. Wijers, M. Jin, M. Wessling, and H. Strathmann, J.Membr. Sci. 147 (1998) 117. Y. Wang and F.M. Doyle, J.Membr. Sci. 159 (1999) 167. X.J. Yang, A.G. Fane, J. Bi, and H.J. Griesser, J.Membr. Sci. 168 (2000) 29. M. Ulbricht, React.Funct.Polym. 31 (1996) 165. T. Yamaguchi, S. Nakao, and S. Kimura, Macromolecules 24 (1991) 5522. T. Yamaguchi, S. Nakao, and S. Kimura, Ind.Eng.Chem.Res. 32 (1993) 848. A.M. Mika, R.F. Childs, J.M. Dickson, B.E. McCarry, and D.R. Gagnon, J.Membr. Sci. 108 (1995) 37. V. Kapur, J. Charkoudian, S.B. Kessler, and J.L. Anderson, Ind.Eng.Chem.Res. 35 (1996) 3179. A.M. Mika, R.F. Childs, J.M. Dickson, B.E. McCarry, and D.R. Gagnon, J.Membr. Sci. 135 (1997) 81. R.F. Childs, A. Pandey, B.E. McCarry, J.M. Dickson, M. West, and J.N.A. Lott, J.Polym.Sci.A:Polym.Chem. 39 (2001) 807. A.M. Mika, R.F. Childs, and J.M. Dickson, PCT/CA97/00770 (WO 98/17377), 1998, Canada. G.E. Gillberg-LaForce and E.M. Gabriel, U.S. Pat. 5,049,275, 1991, U.S.A. A.M. Mika and R.F. Childs, Ind.Eng.Chem.Res. 40 (2001) 1694. D.W. Schaefer, Polymer 25 (1984) 387. J. Happel, AIChE J. 4 (1958) 197. J. Happel, AIChE J. 5 (1959) 174. J.E. Drummond and M.I. Tahir, Int.J.Multiphase Flow 10 (1984) 515. C. McCrory, Impact of Gel Morphology on Pore-Filled Membranes, Ph.D. Thesis, McMaster University, Hamilton, Canada, 2001. A.M. Mika, R.F. Childs, and J.M. Dickson, J.Membr. Sci., In press. S. Singh, K.C. Khulbe, T. Matsuura, and P. Ramamurthy, J.Membr. Sci. 142 (1998) 111. T. Teorell, Progr.Biophysics Biophys.Chem. 3 (1953) 305. K.H. Meyer and J.F. Sievers, Helv.Chim.Acta 19 (1936) 649. K.H. Meyer and J.F. Sievers, Helv.Chim.Acta 19 (1936) 665. K.H. Meyer and J.F. Sievers, Helv.Chim.Acta 19 (1936) 987. E. Hoffer and O. Kedem, Desalination 2 (1967) 25. T. Tsuru, S. Nakao, and S. Kimura, J.Chem.Eng.Japan 24 (1991) 511. F.A. Morrison, Jr. and J.F. Osterle, J.Chem.Phys. 43 (1965) 2111. R.J. Gross and J.F. Osterle, J.Chem.Phys. 49 (1968) 228. J.C. Fair and J.F. Osterle, J.Chem.Phys. 54 (1971) 3307. G.S. Manning, Acc.Chem.Res. 12 (1979) 443. P. Molyneux, In: Water Soluble Synthetic Polymers: Properties and Behavior, CRC Press, Boca Raton, Florida, 1983, p.1.
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[51]
[52]
R.F. Childs, A.M. Mika, A.K. Pandey, C. McCrory, S. Mouton, and J.M. Dickson, Sep.Purif.Technol. 23 (2001) 507. A.M. Mika and R.F. Childs, J.Membr. Sci. 152 (1999) 129.
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Sensors
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New Insights into Membrane Science and Technology:Polymeric and Biofunctional Membranes D. Bhattachatyyaand D.A. Butterfield (Editors) © 2003 Elsevier Science B.V. All rights reserved.
Chapter 18
Membranes for the development of biosensors V. G. Gavalas, J. Wang, L. G. Bachas Department of Chemistry and Center of Membrane Science, University of Kentucky, Lexington, Kentucky 40506-0055 1.
INTRODUCTION
According to the Intemational Union of Pure and Applied Chemistry (IUPAC), "a biosensor is a self-contained integrated device which is capable of providing quantitive or semi-quantitive analytical information using a biological recognition element (biochemical receptor) which is in direct contact with a transducer element" [1 ]. The biological reaction takes place in the interface of the sensor with the sample. The recognition of the analyte is based on its interaction with the biological element. This interaction can be either affinity type or of biocatalytic nature (Fig. 1). The efficient transduction of the biological recognition to analytical signal is crucial for the biosensor's successful operation. The selection of the transducer depends mainly on the type of the biological recognition and the changes that it induces on the system.
Biological Recognition
Affinity Antibodies ,
,,
,,
,
Catalysis Enzyme cofactors
RIgA Receptors ,
,
,
,
,,,
Cells Microorganisms ,
,
,
,
Transduction of biological recognition to signal
Electrochemical llllll
i
ii
i
Thermal i
i
i
Optical i
i
Figure 1. Componentsofbiosensors
379
Mass i
i
i
i i
i
i
i
,,
,,,
,
Membranes For The Development Of Biosensors
2.
-
Bachas
M E M B R A N E S AS COMPONENTS OF BIOSENSORS
Membranes play a key role in the design and development of biosensors. They have been used extensively for the immobilization of biomolecules and/or cofactors, to improve or even alter the selectivity of the biosensor, to control the enzyme kinetics, to improve the biocompatibility of the system, and to control the electron transfer properties of the system. In most cases, the membranes perform more than one of the above functions. These functions are discussed separately in the following sections. 2.1
Membranes for the Immobilization of Biomolecules The biological recognition element in a biosensor is usually immobilized in a layer directly adjacent to or on the surface of the transducer. In most cases, a membrane is used for the immobilization, whether it is a simple physical entrapment or adsorption or more elaborate covalent attachment of the biological recognition elements. In electrochemically mediated immobilization, biomolecules get entrapped within membranes formed by the electropolymerization of appropriate monomers. The final polymeric membrane may be conductive or not, depending on the monomer employed. This is a fast, one-step immobilization procedure, which can be applied essentially to transducers of any shape and size. The thickness of the membrane can be precisely controlled by the charge applied to the deposition of the film. Further details on the electrochemical immobilization of biomolecules can be found in two recently published reviews [2, 3]. For covalent immobilization, enzymes or other recognition elements are tethered to the membranes through chemical bonding between the functional groups of the enzyme and those on the surface of the membrane. However, if more than one such functional groups on the enzyme are available, attachment of the enzyme to the membrane surface through the various groups will inevitably lead to different orientations of the enzyme. The active site of some of the enzyme molecules can become partially or completely blocked by the membrane surface (Fig. 2A). This random immobilization results in reduced enzyme activity due to the limited accessibility of the active site of the enzyme to the sample. The same argument can be made for orienting the binding site of receptors and binding proteins. To minimize activity loss due to random immobilization, the enzyme can be immobilized in an orientation-specific fashion such that the active site points to the same direction that is facing away from the immobilization surface (Fig. 2B). This can be achieved by choosing a unique residue of the enzyme (eg, the SH of a cysteine) through which the protein becomes attached to the membrane. However, choosing such a unique residue may be difficult because (a) most
380
Membranes For The Development O f Biosensors - Bachas
A
Figure 2. Random (A) vs. site-specific (B) protein immobilization. Arrows point to the binding site of the protein.
proteins lack such residues, (b) even if such a residue is present, it may be inaccessible, or (c) the location of the residue may not be ideal for orientation specific immobilization. Recent advancements in molecular biology and other interdisciplinary research efforts have lead to new strategies for orientation specific immobilization of proteins. These strategies include introducing a unique reactive residue into a protein by site-directed mutagenesis, and attaching an affinity tag or a recognition sequence for post-translational modification by gene fusion technology. For example, Huang et al. in our laboratory utilized site-directed mutagenesis to introduce a single cysteine residue to the cysteine-free subtilisin, and this mutant subtilisin was site-specifically immobilized to agarose beads and membranes through the -SH group of the introduced cysteine [4, 5]. Their studies showed that, besides improved stability, an enhancement of activity was observed for the site-specifically immobilized enzyme. Affinity interactions between biomolecules and their ligands have been successfully used in affinity purification of proteins/cells. For example, protein A and protein G have high affinity toward the constant region (Fc) of immonoglobulins (IgG), and therefore, have been used to construct fusion proteins to facilitate purification of other proteins [6, 7]. In essence, the nucleotide sequence coding for the protein of interest is fused to the gene for protein A or protein G. The fused gene is expressed in an appropriate microorganism to get the fusion protein. The fusion protein is readily purified by passing through an affinity column consisting of immobilized IgG. The protein of interest can then be cleaved off the fusion protein and obtained in a purified form for further characterization or application. Other affinity interactions that have been utilized for protein purification include:
381
Membranes For The Development O f Biosensors - Bachas
a. Interactions between binding proteins/domains and their ligand: biotin(strept)avidin [8], cellulose binding domain - cellulose [9, 10], maltose binding domain - amylose [ 11 ], calmodulin - phenothiazine [ 12]. b. Enzymes and their substrates: glutathione-S-transferase - glutathione [ 13], [3galactosidase - p-aminophenyl-13-D-thiogalactoside [ 14]. c. Interactions between affinity tags and their ligands: polyhistidine tag immobilized metal ions (Co 2+, Ni2+), Strep-tag - streptavidin [15, 16], FLAG -antiFLAG antibody [ 17, 18]. d. Others. For example, polyarginine - negatively charged mica surfaces [ 19]. These affinity interactions can be used to help site-specifically immobilize proteins [17, 19]. Such site-specific immobilization enables the active site of the protein to be fully accessible to its substrate even after immobilization. Therefore, high enzyme activity is retained. Site-specific protein immobilization should be advantageous when applied to the design and development of biosensors using immobilized enzymes and other biomolecules. 2.2
Selectivity Control Through Membranes
Interferences in the signal of biosensors can be introduced through two different mechanisms, the biological recognition and the transduction. The first case occurs when the biological element is class-selective (i.e., alcohol oxidase, tyrosinase, etc.) or it can catalyze the reaction of various compounds (i.e., if the biological element is a cell or a tissue). The control of the selectivity in this case is rather difficult since a multitude of compounds can induce a response. When the interference occurs through the transduction pathway, membranes can play a key role in avoiding or minimizing the interference. The most pronounced example of employing membranes to eliminate interference can be found in amperometric biosensors. Since several of the amperometric biosensors are based on the detection of enzymatically produced H202, which requires high potentials (>_ +0.6 V vs Ag/AgC1), electroactive species present in the sample can be oxidized on the electrode surface prohibiting accurate measurements. Even when the applied potential has been lowered by the use of suitable mediators, the signal is still affected since the mediator can promote the oxidation of interferents even at potentials lower than 0.3 V [20-23]. Thus, the use of membranes has been extensively employed to avoid interference from electroactive species such as ascorbic acid, uric acid and acetaminophen [24] by employing one of the following selective permeation mechanisms. Permselective membranes can be based either on size exclusion (cellulose acetate, polyaniline, polypyrrole) or on charge exclusion (Nation, poly(vinylpyridine), poly(ester-sulfonic acid)). The polymeric films used for this purpose are generally solvent-cast or electropolymerized. In general, most of the solvent-cast films are negatively charged and therefore exhibit good rejection only for the negatively charged ascorbic acid 382
Membranes For The Development O f Biosensors - Bachas
and uric acid, but not for the neutral acetaminophen. To overcome this drawback, solvent-cast films based on composite membranes have been employed. Zang et al. [25] used a composite membrane of cellulose acetate and Nation to completely eliminate acetaminophen interference, but the sensor also lost 80% of the H202 sensitivity. Benmakroha et al. [26] demonstrated that by mixing plasticized poly(vinyl chloride) and sulfonated polyether-ether sulfone in suitable proportions, the permeability of hydrogen peroxide over acetaminophen was controlled. Recently, Mizutani et al. [27] used poly-L-lysine and poly(4styrenesulfonate) to form a polyion complex membrane on-top of a gold electrode to prevent electrochemical interferences. The level of selectivity that they obtained was sufficient to allow accurate detection of glucose in blood and beverages. With solvent-cast films it can be difficult to obtain uniform thickness coatings, especially on non-smooth surfaces. In contrast, the thickness of the electropolymerized films can be controlled, and the geometry of the electrode surface does not influence the polymer growth. There are many different monomers that have been used for the fabrication of electropolymerized membranes [2, 3]. Significant impact had the report from Centonze et al. [28] on the use of overoxidized polypyrrole films for the development of interference-free biosensors. The overoxidation of polypyrrole produces a permselective, antifouling film that rejects proteins in addition to the electroactive interferents. It is generally believed that oxygen-containing groups are introduced as a result of overoxidation [29, 30]. Although the detailed identification of these groups has not yet been achieved, these oxygencontaining groups include carbonyl, hydroxyl and carboxyl groups, and they are considered to be responsible for the improved selectivity. Murphy [31] studied a series of electropolymerized films based on derivatives of diaminonaphthalene (DAN) and compared their selectivity and stability with electropolymerized films prepared from o-, m-, or pphenylenediamine. His experiments showed that the positioning of the amino groups greatly influences the selectivity and stability of the sensor. The best results were obtained when 2,3-DAN, 1,5-DAN or 5-amino-l-napthol were employed. 2.3
Controlling enzyme kinetics
The kinetics of many enzymes can be described with the model that Leonor Michaelis and Maud Menten have proposed [32]" E+S
~
ES
It,. AL~
~ r
p
k_l
383
Membranes For The Development O f Biosensors - Baehas
The enzyme (E) reacts with the substrate (S) to form a complex (ES) with rate constant kl. The complex ES can either dissociate into E and S (with rate constant k_l) or form the product P (with rate constant k2). Introducing the Michaelis constant (KM) as the ratio (k.1 + k2)/kl and assuming that (a) the ES complex is in steady state, (b) under saturating conditions all the enzyme is converted to ES complex, and (c) the maximum velocity (Vmax) is achieved when all the active centers of all the enzyme molecules are saturated with substrate, then the velocity (v) is defined by the equation: V = Wmax
IS] [s] + KM
(1)
From this non-linear equation it is clear that the velocity, v, will have a linear relationship with the substrate concentration, if this concentration is low relative to KM (i.e. when KM >_0.1 IS]). The KM of most enzymes is in the low mM range (i.e., glucose oxidase has a KM value of 4.2 mM for glucose, L-lactate oxidase a KM of 1.4 mM for lactic acid, L-lactate dehydrogenase a KM of 1.6 mM for lactic acid, malate dehydrogenase a KM of 0.2 mM for maleic acid, etc.), while the concentration of the corresponding analytes in real samples is much higher (glucose up to 500 mM in food samples,-~5 mM in blood; lactate can exceed 7 mM in blood; malate concentrations are greater than 5 mM in food). Thus, in order to increase the linear range of response, membranes that limit the substrate diffusion can be used. Numerous models have been reported that describe the effect of diffusion processes in the response of enzyme biosensors of various configurations [3337]. Cambiaso et al. [38] calculated that decreasing the substrate diffusion coefficient from 1.1xl0 9 cm2/s to 5x10 -1° cm2/s, the KM is increased by 50%. Schulmeister and Pfeiffer [39] developed a model to describe the effect of using a semipermable membrane of variable thickness on the dynamic behavior of a glucose biosensor. In Fig. 3 the correlation of the linear range of response of the glucose biosensor with the thickness of the additional perforated membrane is presented. As it can be seen, by using a thicker membrane the linear range of response is extended, and this extension is obtained because the enzymatic process is controlled by the diffusion of the enzyme's substrate. However, by limiting the diffusion of the substrate, the sensitivity of the biosensor is decreased as well. Schulmeister and Pfeiffer [39] calculated that by increasing the membrane thickness from 5 to 40 ~tm the sensitivity of the biosensor is reduced by 85%. Thus, there is a trade off between linear range and sensitivity, which has to be considered when designing a biosensor.
384
Membranes For The Development Of Biosensors - Bachas
50 E "E m ~t ,m .--I
40
._. 30 u
~: w 20
7 o
~
;o
~
~'o
~
~o
~
4o
,~
Membrane Thickness (pm)
Figure 3. Correlation of the effective linearity of a glucose biosensor with the thickness of the semipermable membrane. Data from Schulmeister and Pfeiffer [39] Various membranes have been used as extemal diffusion control layers, including cellulose acetate/triacetate, polycarbonate, polyurethane, Nation, polyvinylchloride (PVC), etc. Yang et al. [40] used a PVC coating to increase the upper range of linear response of glucose and lactate biosensors, from 1 to 24 mM and from 2 to 16 mM, respectively. The sensitivities of the PVC-coated sensors were similar, despite the significant difference in the activities of the enzymes. Upon the use of the PVC membrane, the contribution of the enzyme activity to the sensor's signal is minimized, since the membrane provides sufficient diffusion limitation for both substrates. In a recent systematic study, Maines et al. [41] evaluated a series of modified membranes on the basis of physical and diffusion properties, by calculating the diffusion coefficients of the enzyme's substrate and various interferents. Further modification of the membranes allowed for the optimization of additional parameters (i.e., stability and selectivity) other than the linear range of response.
2.4
Membranes for Improving the Biocompatibility of Sensors In vivo monitoring of various analytes is important for many bioanalytical
and biomedical applications. The crucial challenge in this type of application is the interactions of the sensor with the host environment, which are qualitatively described by the term biocompatibility. The real issue of sensor biocompatibility should not only be whether there are adverse biological reactions to the sensor, 385
Membranes For The Development O f Biosensors - Bachas
but whether the sensor has any negative effect on the body. The term biocompatible should suggest that the material described exhibits harmonious behavior when in contact with tissue and body fluids [42]. However, biocompatibility has been used occasionally in bioanalytical chemistry to describe the effect of the body on the sensor. Insertion of a biosensor in the body of humans or animals may stimulate various "foreign body response" processes, which can lead to the failure of the sensor. The most important reactions that follow the contact of a sensor with blood are protein adsorption, platelet adhesion, platelet aggregation, fibrin formation and thrombus formation [43]. Biodegradation of the membrane, breakdown/denaturation of the biological recognition element (e.g., the enzyme) or other components of the biosensor could also lead to its malfunction. The most common failing reasons of sensors in vivo are presented in Fig. 4.
Membrane biodegradation
Biofouling
Degradation of recognition element
Fibrous encapsulation
Figure 4. Schematic illustration of failing reasons ofbiosensor in vivo.
Several recent reviews discuss methods and materials used for the improvement of sensor biocompatibility [44-46]. One of the most important goals in the design of biocompatible membranes is to develop materials that resist biofouling. To avoid biofouling, the sensor's surfaces have been treated with anticoagulants (i.e., heparin) or surfactants. Brooks et al. [47] have covalently attached heparin on the surface of cellulose triacetate membranes without any significant alteration on the response characteristics of the sensor (potassium ISE based on valinomycin). The cellulose triacetate membranes, which were coated with heparin, inactivate blood coagulation factor Xa, increasing the blood compatibility of the sensor. Reddy and Vadgama [48] compared the performance of a series of commercial membranes with PVC membranes that have been modified with nonionic (Tween 82 or Triton X-100), anionic (bis(2-ethylhexyl)hydrogenphosphate) or cationic surfactants (tricaprylylmethyl-ammonium chloride). The cationic surfactant-modified PVC membrane exhibited the worst biocompatibility. The other modified PVC membranes had improved biocompatibility compared to commercial porous
386
Membranes For The DevelopmentOf Biosensors- Bachas
polycarbonate and Cuprophan hemodialysis membranes. However, a lactate biosensor lost between 35 to 80% of its response upon 6 h exposure to blood, regardless the type of membrane used. Berrocal et al. [49] introduced the use of Biospan-S, a new bloodcompatible polymeric material (Fig. 5A). Biospan-S is segmented polyurethane modified with poly(dimethylsiloxane) and exhibits higher in vivo stability than the unmodified segmented polyurethane [50]. The authors prepared three different types of cation-selective electrodes, which demonstrated comparable analytical characteristics to conventional PVC-based electrodes, verifying that the Biospan-S can be used as a polymeric matrix for the development of bloodcompatible sensors. Recently, new polymeric materials that mimic the phospholipid groups naturally occurring in membranes of biological cells have been used (Fig. 5B). Berrocal et al. [51] used poly(2-methacryloyloxyethyl phosphorylcholine-cobutylmethacrylate), a biocompatible hydrogel containing phosphorylcholine groups, as a coating on the membrane surface of ion-selective electrodes to improve their compatibility. They demonstrated that the hydrogel is compatible with several membrane materials, introducing slight alterations in the selectivity but leaving unaffected the sensitivity and detection limit of the sensors. In addition, immunostaining experiments verified that there was limited platelet adhesion to the coated PVC membranes. This hydrogel has also been used to improve the biocompatibility of a glucose biosensor [52].
CHa-Si-Ol"-'Si-OcH3 LCH3 Jp
(A)
"
'
-('CH,.CH,CH,CH,O),.,.,
-0-"
hI'-o-"
o-.-s,-lo-s,-ro-s,-c.
"
z ,
CH3 (B)
c. 3[
ca3Jp
Ca3
CH3
( I--C2 )a ( ? O-'-C O O~ I , + I O C H 2CH20~OCH 2CH2N(CH3)3 ocn
2 c n 2CH 2 c n 3
O"
Figure 5. Structure of (A) Biospan-S and (B) poly(2-methacryloyloxyethyl phosphorylcholine-co-butylmethacrylate) or poly(MPC-co-BMA) 387
Membranes For The Development O f Biosensors - Bachas
Another approach is the use of materials that continuously release biological active molecules such as nitric oxide to prevent platelet adhesion and activation. Schoenfisch et al. [53] prepared a silicone rubber membrane containing NO-generating diazeniumdiolate compounds and applied it on the outer surface of an oxygen catheter. The NO-generating diazeniumdiolate compounds, when in contact with water, release NO for a substantial period of time, without altering the analytical characteristics of the sensor. The performance of the membrane-covered sensor was evaluated in vivo in dogs with very promising results. Despite the extensive research, the issue of sensor biocompatibility remains a significant challenge. More fundamental research on the sensor-tissue interactions will be needed to fully understand and eventually overcome the problems related to membrane biofouling. 2.5
Membranes for Electron Transfer Control
The function of amperometric biosensors is based on the electron transfer from the active site of the immobilized enzymes to the electrode surface. The enzyme during the catalysis stores the electrons either in a prosthetic group (i.e., FAD, PQQ, heme, etc.) or in a co-substrate (i.e., NAD(P)+). The amperometric biosensor's signal is based on the transfer of these electrons from the intermediate storage group to the electrode surface [54]. It is, thus, evident that the analytical characteristics of a biosensor can be optimized through the control of the electron-transfer process. Membranes that contain redox active centers, known as "redox polymers", have been extensively used for the development ofbiosensors [54, 55]. In Fig. 6 the electron pathway in a redox polymer-based biosensor is shown. The electrons, through an electron-hopping mechanism, are transferred from the enzyme to the electrode surface. One of the most successful examples of redox polymer is the osmium based poly(vinyl imidazole) polymer introduced by Heller's group [56, 57]. Glucose oxidase has been "wired" (the enzyme was connected to the electrode through cross-linked electron-conducting redox hydrogel) with polyelectrolytes having electron relaying redox centers [Os(bpy)2C1]+/2+ in their backbones. Hydrogels were formed upon cross-linking the enzyme with the polymer. The detection was achieved by applying +0.3 V potential to the electrode. Osmiumbased polymers have been used to promote electron transfer from other enzymes as well; e.g., horseradish peroxidase [58-60], lactate oxidase [61, 62], glutathione sulfhydryl oxidase [63]. Swann et al. [64] reported on the use of the conductive polypyrrole polymer as the electron transfer pathway. They observed electron transfer from the PQQ of fructose dehydrogenase to the electrode, which was affected by the polypyrrole growth conditions. More widespread is the use of redox-modified 388
Membranes For The Development O f Biosensors - Bachas
monomers for the growth of polymeric film. There have been reports on osmium-complex-modified pyrrole [65-67], ferrocene-modified pyrrole [68], or aniline [69-71 ]. The use of such redox-polymers has been attractive since it involves a one-step preparation procedure that enables the control of the polymer thickness even when the electrode surface is rough, small sized, or highly porous.
Figure 6. Electron-hopping mechanism in a redox polymer-based biosensor (-- polymer backbone, • redox active center). Arrows denote possible paths of electron transfer. 3.
CONCLUSIONS AND FUTURE CONSIDERATIONS
Membranes play an indispensable role in the development of biosensors, not only because most electrodes surfaces themselves are membranes, but also because more and more polymeric membranes are developed/functionalized to improve the selectivity, sensitivity, and operational range of biosensors. Since sensor performance characteristics are largely determined by the biological recognition element of a biosensor, increasing efforts are devoted on how to gain a well-controlled and reproducible immobilization of those biomolecules. Though relatively simple, entrapment of biomolecules into polymeric films through electrochemical polymerization has the drawbacks of reduced biological activity. Site-specific immobilization, in which all the active sites of the immobilized biomolecules point to the direction toward the solution, provides a new way to a well-controlled sensing layer with high biological activity. The
389
Membranes For The Development O f Biosensors - Bachas
drawback of the initial investment in modifying the biomolecules either chemically or genetically is outweighed by the gain of highly active immobilized biomolecules on the sensor surface. The need for smaller biosensors with faster response as well as array sensors for multi-analyte sensing makes the preparation of the sensor surfaces more challenging. Recent advancement in micro-patterning of biomolecules and cells may provide useful tools to gain well-defined immobilization of biomolecules with superior activity at a molecular scale. For example, self-assembled monolayers (SAMs) using microcontact printing allow patterned formation of (bio)molecules and cells on gold and silicon-based membrane surfaces [72, 73]. This technique has been recently extended to include microcontact printing on polymeric materials such as poly(ethylene teraphthalate) [74]. These techniques can be used to mass produce sensor surfaces in the micrometer scale with delicate biomolecules at a reasonable cost. Polymeric microcapsule arrays have been synthesized by Martin and coworkers, and subsequently used as cavities to immobilize enzymes [75, 76]. Recently, the same group has successfully produced membranes formed by orderly aligning nanotubules modified with HS-R, which can selectively transport hydrophobic/hydrophilic species depending on the nature of the R group [77, 78]. These nanotubules should be suitable for the immobilization of small enzymes and other biomolecules and lead to the development of biosensors at the nanoscale.
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392
New Insights into Membrane Science and Technology: Polymeric and Biofunctional Membranes D. Bhattacharyya and D.A. Butterfield (Editors) © 2003 Elsevier Science B.V. All rights reserved.
Chapter 19
Ion-partitioning membranes as electroactive elements for the development of a novel cation-selective CHEMFET sensor system E. A. Moschou .1, N. A. Chaniotakis* Laboratory of Analytical Chemistry, Department of Chemistry, University of Crete, 71409 Iraklion, Crete, Greece 1Current address" Department of Chemistry, University of Kentucky, Lexington KY 40506-0055 1.
ABSTRACT
The development of a cation-selective sensor based on a new response mechanism is presented in this paper. The sensor is based on a bulk active ionpartitioning membrane as the chemical recognition element while the signal transduction is obtained by a pH sensitive transducer, such as the pH-ISFET. The ion-partitioning membrane is doped with two different ionophores, one selective to the analyte cation and the other to protons. The increase of the analyte cation activity in the sample results in the carrier mediated partition of the analyte into the bulk of the membrane phase. Due to the electroneutrality principle requirements, a simultaneous displacement of protons of equal charge out of the membrane must take place. This proton flux takes place towards both the membrane/sample and membrane/transducer interfaces. The monitoring of the proton flow in or out of the membrane by the pH-sensitive transducer is directly related to the analyte ion activity in the sample. Based on this mechanism, it is shown that there is a large effect of all membrane components, as well as the ionic composition of the sample, on the sensor selectivity, sensitivity and detection limit. It is also shown that the optimization of the primary cartier to the proton ionophore ratio in the membrane and that of the properties of the proton ionophore are vital in obtaining the desired sensor characteristics. It is shown that the use of a plasticizer and a polymer matrix with low dielectric constant favors the cationprotons exchange mechanism stabilizing the observed signal. Finally, the proper adjustment of the sample pH and the use of hydrophilic anionic counteranions in the sample contribute to the improvement of both the sensitivity and the detection limit of the sensor. The analytical characteristics of the CHEMFETs selective to the monovalent potassium ion and the divalent calcium ion with the 393
Ion-Partitioning Membranes As Electroactive Elements For The Development Of A Novel Cation-Selective CHEMFET Sensor System Moschou
optimized ion-partitioning membranes are evaluated by the analysis of blood serum samples. 2.
INTRODUCTION
Polymeric membranes have been extensively used as the chemical recognition element in the design of chemical sensors. They are water immiscible phases of high viscosity and are based on the polymer matrix, which plays the role of the supporting material; and the plasticizer, which serves as the organic solvent of the membrane phase. The chemical recognition of the analyte ion, Iz, is based on the partition of the hydrophilic ion from the aqueous sample into the lipophilic membrane phase. This partition is expressed by the partition coefficient of the ion, ki, and is affected by the dielectric constant of the medium, and the size and charge of the analyte ion [1,2]. The increase of the ion size and the decrease of the ion charge result in the increase of the ion partition in the lipophilic membrane. An experimental partitioning series of the ions according to their lipophilicity in the octanol-water system has been formulated, and is the so-called Hofmeister selectivity series [3]: Cs ÷ > Rb ÷ > K ÷ > NH4÷ > H ÷ > Na + > Li ÷ > Ba 2+ > Sr2÷ > Ca 2+ > Mg 2÷ (1) C104" >
SCN" > I" > NO3- > Br > NO2- > C1- > HCO3
>
5042- > F-
(2)
In order to deviate from the Hofmeister selectivity series and increase the relatively low partition of the hydrophilic ions into the lipophilic membrane phase, the ionophores are added in the polymeric membrane. These ionophores [4,5] are mobile electrically neutral or charged, lipophilic, organic or inorganic substances that can selectively interact [6,7] with the analyte ions through electrostatic interactions, wan der Waals or hydrogen bonding. This interaction facilitates the transport of the analyte ions into the lipophilic polymeric membrane. In addition, the polymeric membranes are also doped with electrically charged, organic substances called lipophilic ionic sites [8,9]. These charged compounds decrease the electrical resistance of the membrane, provide charged sites and labile counterions, and can thus increase the partitioning of the analyte ions in the lipophilic polymeric membrane. The response mechanism of the polymeric membrane as a chemical recognition element is determined by its chemical composition. If the membrane is doped with one ionophore that is selective to the analyte ion of interest, this membrane is called surface active membrane [10]. There is a direct chemical interaction between the ionophore and the analyte ion. This interaction mediates
394
Ion-Partitioning Membranes As Electroactive Elements For The Development Of A Novel Cation-Selective CHEMFET Sensor System Moschou
the ion transport from the aqueous sample to the outer phase boundary of the membrane. The ion extraction is restricted to the membrane/aqueous solution Surface Active Polymeric Membrane AqueousSample L
R-
LI + +
I+ A-
LI +
I+ A"
L : Electrically neutral ionophore
L-
R+
L-: Electrically charged ionophore Figure 1. Response mechanism of a surface active membrane; where I+ is the analyte cation; A is the counteranion in the aqueous sample; L and L- stand for the electrically neutral and charged ionophore, respectively; and R, R+ are the lipophilic anionic and cationic sites, respectively. interface due to the electroneutrality principle. Therefore, according to the surface active membrane response mechanism (Figure 1) the analyte ions are extracted into the outer membrane phase boundary to a final concentration, which is related to the activity of the analyte ion in the aqueous sample. The one-carrier-doped surface active polymeric membrane can be further modified with the addition of a second ionophore, to produce the so-called bulk active membrane [10]. This ionophore can be selective to either an ion of opposite charged, or of the same charge, which is usually the proton. The increase of the analyte ion activity in the sample results in the ionophoremediated transport of the ion from the aqueous sample to the polymeric membrane. The charge accumulation alteration due to the primary ion partitioning into the membrane can now be compensated by the extraction of the secondary ion of analogous charge out of the membrane for electroneutrality reasons. For a specific application when the analyte ion is a cation, its partitioning into the polymeric membrane will promote the extraction of protons of equal charge out of the membrane. Due to the ion-exchange of the analyte cation with protons the response mechanism of this kind of bulk active membranes is called ion-exchange mechanism [11,12] (Figure 2). In the case of anionic analytes, protons of opposite charge must be co-extracted into the membrane to fulfill the electroneutrality requirements. Due to the simultaneous 395
Ion-Partitioning Membranes As Electroactive Elements For The Development Of A Novel Cation-Selective CHEMFET Sensor System Moschou
Bulk Active Membranes Based the Ion-Exchange Mechanism
Based on an electrically neutral ionophore L /
I+ ~
LI +
L
I+ ~1~ LI +
L
H+
C H + R-
C
H + : = ~ CH
C"
Based on an electrically charged ionophore L/
/
I+ ; = ~ LI H + m : ~ CH
R+
I+ ~
L" C"
LI
L"
H + ~=~ CH + C
I+ ~
LI
H+
LH
Figure 2. Bulk active membranes based on the ion-exchange mechanism. [I÷ is the analyte cation; L, LI+ stand for the free and complexed ionophore, respectively; C, CH + are the unprotonated and protonated ionophore and R+, R the lipophilic ionic sites].
Bulk Active Membranes Based on the Co-Extraction Mechanism
Based on electrically neutral ionophore L /
X:;~ H+~ : ~
LX CH
R+ L C
X••
LX"
L
CH + C
Based on electrically charged ionophore L+ /
X"
LX
H÷
CH +
R"
L+
X" ~
LX
L+
C
H ÷~=~ CH
C"
Figure 3. Bulk active membranes based on the co-extraction mechanism. [X is the analyte anion; L, LX" stand for the free and complexed ionophore, respectively; C-, CH are the unprotonated and protonated ionophore and R-, R+ the lipophilic anionic and cationic sites, respectively].
396
Ion-Partitioning Membranes As Electroactive Elements For The Development Of A Novel Cation-Selective CHEMFET Sensor System Moschou
co-extraction of anions and protons in the bulk of the polymeric membrane, these active membranes operate based on the so-called co-extraction mechanism [13,14] (Figure 3). In conclusion, in the case of the surface active membranes, the analyte ions are partitioned into the outer phase boundary of the membrane generating a potential difference [15]. For this reason this type of membranes are used as the chemical recognition element for the construction of potentiometric sensors. In contrast, in bulk active membranes, the ions of interest are partitioned into the bulk of the membrane phase, with the electroneutrality principle being fulfilled due to the simultaneous partition of the secondary ions [ 16]. Therefore, the bulk active membranes cannot be used as the chemical recognition element for the construction of potentiometric sensors, but they are rather used in optical sensors (i.e. the ionophore selective to the secondary ion is a chromoionophore and is changing its optical properties upon the analyte ion/secondary ion partition in the membrane). The most significant example of the optical sensors based on bulk active polymeric membranes are the optodes [ 17,18], while the main representatives of potentiometric sensors based on surface active polymeric membranes are the ion selective electrodes [19,20] (ISEs) and the chemically modified field-effect transistors [21,22] (CHEMFETs). In the ISEs the internal reference electrode (usually a Ag/AgC1 electrode) and the internal filling solution (containing a known activity of the analyte ion) are used for the signal transduction. In the CHEMFETs the reference electrode and the internal filling solution are substituted [23] by a pH sensitive field-effect transistor (pH-ISFET) [24,25]. This design allows for the elimination of the internal filling solution. For this reason CHEMFETs are attractive sensor systems since they can be easily miniaturized [26,27]. The CHEMFETs are thus constructed (Figure 4) ivy the application of an electroactive polymeric membrane onto the gate of the pH-ISFET. When the sensor is immersed in the sample solution, the target ions are partitioned into the outer phase boundary of the polymeric membrane developing a potential difference [28,29]. This potential difference is measured via field-effect of the pH-ISFET generating the sensor signal. The pH-sensitive metal oxide [30,31] gate of the pH-ISFET is thus active only in the case of a pH-sensitive sensor system [32]. Its electrochemical activity is subdued when the polymeric membrane is applied on the pH-ISFET's gate. The metal oxide gate does not participate in the mechanism of the signal generation of the CHEMFET any longer [33,34]. Recently it has been shown in the literature that water molecules diffuse into the bulk of the polymeric membrane and can thus form a thin aqueous layer at the interface between the polymeric membrane and the gate of the pH-ISFET 397
Ion-Partitioning Membranes As Electroactive Elements For The Development Of A Novel Cation-Selective CHEMFET Sensor System Moschou
[35,36]. When the sensor is immersed in a sample containing volatile acids or
Reference Electrode
r"!
[[
PolymericMembrane
U
@1
; ~alyte ~aiion
'AIOH~
I S°urce I
= = = I Drain substrate Si
I
T Figure 4. Schematic representation of a CHEMFET based on a polymeric membrane. bases, these compounds can diffuse through the polymeric membrane and dissociate in the thin aqueous layer [37]. This dissociation alters the pH at the gate-membrane aqueous layer, producing a signal transduced [6,38]. This was the first report of the pH-ISFET's gate playing an active role on the signal generation of the CHEMFET even though the response to these species was considered as interference. This interference has been overcome by the introduction of a pH-buffered layer between the pH-ISFET's gate and the polymeric membrane [39,40]. This way the gate of the pH-ISFET is in contact with a constant pH environment, and is thus passivated. In this paper we present the development of a cationic electrochemical sensor based on a bulk active ion-partitioning membrane as the chemical recognition element while the signal transduction is obtained by a pH-ISFET. The ion-partitioning membrane is doped with the ionophore selective to the analyte cation as well as an ionophore selective to the proton as the secondary ion. Due to the chemical composition of the polymeric membrane, the ionexchange of the analyte cations with protons takes place in the bulk of the membrane. The displaced protons exit the membrane from both the outer and the inner phase boundaries towards the aqueous sample and the pH-ISFET's gate, respectively. The proton transducer monitors the membrane proton flux generating the CHEMFET's signal [41,42]. Therefore, the direct measurement of the membrane proton flow by the pH-ISFET is translated into the indirect determination of the analyte cation activity in the sample. The potentiometric response of the CHEMFET based on the ion-partitioning membrane is based on an original response mechanism where the chemical recognition element
398
Ion-Partitioning Membranes As Electroactive Elements For The Development Of A Novel Cation-Selective CHEMFET Sensor System Moschou
interacts actively with the signal transducer. In addition, due to the direct determination of the displaced membrane protons by the pH-ISFET, the actual sensitivity of the sensor is 59.2mV/pH for every cation analyte determined, and is translated into the sensitivity of 59.2mV/[IZ÷]. Therefore, the main advantage of the CHEMFET based on an ion-partitioning membrane is the fact that its sensitivity is 59.2 mV/decade of the analyte cation activity and therefore is not proportional to the inverse of the analyte's charge as for the conventional potentiometric sensors. As a result the CHEMFETs based on an ion-partitioning membrane present the super-Nemstian sensitivity of 59.2 mV/decade [I"+] for the multivalent cations Iz+. In this paper we present the development of CHEMFETs based on ionpartitioning membranes that are selective towards monovalent and divalent cations. Potassium and Calcium were selected as the model cations. The effect of the composition of both the sample and the ion-partitioning membrane is evaluated as a function of the magnitude of the membrane proton flux, and thus the observed sensitivity of the resulting CHEMFETs. The analytical characteristics of the CHEMFETs based on the optimized ion-partitioning membranes are evaluated with the determination of the free potassium and calcium ion activity in blood serum samples. 3.
EXPERIMENTAL
3.1
Reagents and Membranes All reagents used were of puriss p.a. grade and were obtained from Fluka, while deionized water (Bamstead NAN-O-Pure) was used for the preparation of all solutions. All the polymeric membrane components were Selectophores and were purchased from Fluka. The proton carriers used were ETH 1907, 4nonadecylpyridine, ETH 2439, [9-Dimethylamino-5-[4-(16-butyl-2,14-dioxo3,5-dioxaeicosyl)phenylamino] benzo[a] phenoxazine and tridodecylamine (TDA). Valinomycin was incorporated as the K÷ ionophore and potassium tetrakis (4-chlorophenyl)borate served as the lipophilic anionic sites for the potassium selective ion-partitioning membranes. The calcium ionophore used in the calcium selective ion-partitioning membranes was the ETH 1001 [((-)-(R,R)N,N'-(Bis( 11-ethoxyc arbonyl)undecyl-N,N'-4,5-tetramethyl-3,6dioxaoctanediamide with the lipophilic anionic sites potassium tetrakis [3,5bis(trifluoromethyl)phenyl] borate. The plasticizers tested were bis(2ethylhexyl) sebacate (DOS), 2-nitrophenyl octyl ether (o-NPOE) and tributyl phosphate (TBP). The polymer matrixes examined were high molecular weight poly(vinyl chloride) (PVC), carboxylated PVC (PVC-COOH, kindly provided by Prof. Mark E. Meyerhoff, Department of Chemistry, University of Michigan, Ann Arbor), aminated PVC (PVC-NH2) and silicone rubber (constructed by the
399
Ion-Partitioning Membranes As Electroactive Elements For The Development Of A Novel Cation-Selective CHEMFET Sensor System Moschou
combination of siloprene K1000 and siloprene crosslinking agent K-11 in CH2C12 directly before membrane application). The ion-partitioning membranes constructed contained valinomycin or ETH 1001 (1%wt), the proton carrier ETH 1907 at a molar ratio 1 unless otherwise stated, the lipophilic anionic sites (0.3 %wt), the plasticizer and polymer matrix (in 2:1 weight ratio) in either THF (for PVC and substituted PVC-based membranes) or CH2C12 (for the silicone rubber-based membranes) in 1.8 mL solvent per 100 mg of membrane constituents. 3.2
Sensor Construction Aluminum oxide pH-ISFETs (with internal Ag/AgC1 reference electrodes) were provided by Thermo Orion (Catalog No. 615700). The CHEMFETs were prepared by casting 8 ~tL portions of the ion-partitioning membrane cocktail onto the pH-ISFET's gate and were allowed to dry at room temperature for 30 minutes (estimated membrane thickness less than 10 ~tm). The optimized Ca-CHEMFETs incorporated a dextran layer between the pHISFET's gate and the ion-partitioning membrane. The role of this dextran layer was to stabilize the thin aqueous layer formed between the ion-partitioning membrane and the pH-ISFET's gate, where the membrane proton flow is measured. These Ca-CHEMFETs were constructed by casting a small amount (0.3~tl) of the diethylaminoethyl dextran (Sigma, D-9885) solution prepared by dissolving 10mg dextran in 0.5 mL deionized water, on the pH-ISFET's gate. After 90 minutes, the cocktail of the calcium ion-partitioning membrane was deposited over the dextran layer. The CHEMFETs were then soaked in the testing solution for 30 minutes. Solutions with 10~-10 -1 M KC1 and low initial pH values adjusted with 1 M HC1 were used for the pH-response studies. The pH was increased from low to high pH values using either 1 M Tris (Tris(hydroxymethyl)aminomethane) or 2.5 M NaOH. The calibration curves were made by standard additions of 1 M KC1 or CaC12 to 10-5 M solutions of the respective cation buffered with 0.01M Tris-HC1 to pH 7.5. 3.3 Electrochemical Measurements
The CHEMFET signal was transduced to potential changes by a model 605 preamplifier (Thermo Orion). A Thermo Orion ROSS pH electrode was used to monitor the sample pH during all experiments while the laboratory temperature was set to 25+0.5 ° C. The data were collected using a personal computer with software written in Basic.
400
Ion-Partitioning Membranes As Electroactive Elements For The Development Of A Novel Cation-SelectiveCHEMFET Sensor System Moschou
3.4
Blood Serum Analysis The flee potassium and calcium ion activity in blood serum samples was measured by the K-CHEMFETs and also a potassium [43] and calcium [44] selective ISE as the control method [45,46]. Both of the above sensors were inserted in a flow injection analysis (FIA) system. The FIA system consisted of one pulseless pump (Sage 362, Thermo Orion), and two units, the unit of the potassium determination with the KCHEMFET and the K-ISE in series and the second unit in parallel were the CaCHEMFET was followed in series by the Ca-ISE for the calcium measurement. So, the system was constructed by the use of two V-100D diagonal flow injection valves (Upchurch Scientific) with 230 gL loop volumes and two reaction coils (0.030 in. i.d., 100 cm length). The four potentiometric flow through cells used were constructed of Derlin (Goodfellow Cambridge Ltd.) and were used in series (the outlet of the K-CHEMFET flow through cell was connected to the inlet of the K-ISE flow through cell and in analogous way the Ca-CHEMFET and Ca-ISE were connected to each other). The stream carriers used was 10-2 M Tris-HC1 pH 7.5 for the potassium determination and a 102M potassium hydrogen phthalate (KHP) pH 3.0 for the calcium measurement under a flow rate of 2.5 mL/min. The sensors were calibrated by the use of artificial blood serum solutions containing the following electrolytes [47]: 5.0x10 "4 M NaHSO4, 8.5x10 "4 M MgCI2, 3.0xl 0 .2 M NaHCO3, 9.5xl 0 .4 M NaH2PO4, 1.4xl 0 -1 M NaC1 and either 2.5x10 3 M CaCI2 or 4.0x103M KC1 for the potassium and calcium standards, respectively. The blood serum samples and the artificial blood serum solutions were diluted five times (ratio 1:5) using the buffer Tris-HC1 pH 7.5; while the blood serum sample was either diluted five to eleven times (ratios 1:5 up to 1"11) for the potassium determination or standard calcium additions took place for the calcium measurement to provide blood serum samples of various potassium and calcium concentrations in the physiological range. 4.
RESULTS AND DISCUSSION
Due to the composition of the membranes employed, any increase of the analyte cation activity in the sample will result in the mediated transport of the analyte cations into the bulk of the polymeric membrane with the simultaneous transport out of the membrane of the secondary cation already present in the membrane.
I
z+
(aq) +
nL(org) +
zCH+(org)
<--)'
L nXTz+(org) +
401
zC(org) +
zH+(~a)
(3)
Ion-Partitioning Membranes As Electroactive Elements For The Development Of A Novel Cation-Selective CHEMFET Sensor System Moschou
where L, LnIz+ is the free and eomplexed ionophore selective to the analyte cation Iz+ and C, CH + is the unprotonated and protonated H + ionophore. The displaced protons leave the membrane by both the outer phase boundary towards the aqueous solution and the inner phase boundary of the membrane towards the pH-ISFET's gate. The pH-sensitive transducer monitors the membrane proton flow generating the CHEMFET's signal. The magnitude of the sensor signal is affected by the magnitude of the membrane proton flow, which in turn is governed by the ion-exchange equilibrium constant, KLnXexeh: KLnlexch =
([LnIz+] / ai" [L]n) "(all " [C] / [CH+])z= (Ka / kH)Z "kI " ~LnI
(4)
where K a is the dissociation constant of the proton ionophore, ~LnI the analyte cation-ionophore formation constant and ki, kH the partition coefficients of analyte ions and protons, respectively. The brackets imply the concentration of the species in the membrane phase. In addition the ion-partitioning is also governed by the proton transport process, which contributes to the magnitude of the total membrane proton flow. The proton ionophore is able to extract and mediate the transport of protons from the aqueous sample through the bulk of the membrane to the inner membrane boundary, according to the following equation [48,49]:
(5)
H+(org) = [~CH " [C] " kH " aH
where H+(org)is the proton concentration partitioned in the membrane phase, ~CH is the complex formation constant of the proton ionophore/H + complex, [C] the concentration of the proton ionophore in the membrane, kH the partition coefficient of protons and an the proton activity in the aqueous sample. From equations 3, 4 and 5 it can be seen that the membrane proton flow governing the CHEMFET's signal is affected by the composition of the aqueous sample and the ion-partitioning membrane constituents. In order to evaluate the effect of these parameters on the potentiometric response of the CHEMFET the potassium and the calcium ions are chosen as the model monovalent and multivalent cation, respectively.
4.1
Effect of the Chemical Composition of the Aqueous Sample According to equation 3, the increase of the potassium ion activity in the sample will result in the increase of the amount of protons displacing out of the ion-partitioning membrane. In addition, according to equation 5, the decrease of the sample pH will increase the magnitude of the proton transport through the bulk of the membrane. These two processes, even though they are competitive, they both contribute to the membrane proton flow measured by the pH-ISFET.
402
Ion-Partitioning Membranes As Electroactive Elements For The Development Of A Novel Cation-Selective CHEMFET Sensor System Mosehou
200 150 100 50
1.0E-02 1.0E-04 n 2.5
4.0
5.5
7.0
8.5
10.0
~'~
pH Figure 5. Effect of the proton and potassium ion activity on the potentiometric response of a potassium selective CHEMFET.
In samples of low pH value the proton transport process dominates on the development of the sensor's response (Figure 5). On the contrary, in samples of high pH, the ion-exchange equilibrium of K + and H + is the one determining the signal generation of the CHEMFET. In the region in between, both processes are significant and affect the potentiometric response of the sensor.
4.2 Effect of Ion Exchange Equilibrium In addition to, the ion-exchange equilibrium is also influenced by the composition of the ion-partitioning membrane used (Eq. 2 and 3). The relative concentration of the analyte and proton ionophores incorporated in the membrane affect the ion-exchange equilibrium favoring either the analyte partitioning or the proton transport within the bulk of the membrane. As indicated by Eq. (5), the decrease of the proton ionophore concentration in the membrane hinders the proton transport from the aqueous sample through the membrane to the pH-ISFET's gate. For this reason the increase of the molar ratio of the potassium to proton ionophore decreases dramatically the linear range of the CHEMFET's response to pH by (Figure 6). On the contrary, the increase of the K + ionophore relative to the H + counterpart results in the increase of the potassium sensitivity of the sensor as it increases from 45 to 55 mV/[K +] for a molar ratio increase from 0.5 to 1.0. This increase is due to the enhanced of K ÷ to H ion-exchange in the membrane phase (as dictated by Eq. 4) that
403
Ion-Partitioning Membranes As Electroactive Elements For The DevelopmentOf A Novel Cation-Selective CHEMFET Sensor System Moschou
generates the increased membrane proton flow measured by the pH-sensitive transducer.
250 200 ~" 150 100 2.0 .0 5
50
2.5
4.0
5.5
7.0
8.5
10.0
pH Figure 6. Effect of the molar ratio of the K+ to H+ membrane ionophores on the potentiometric response of the K-CHEMFET varying the pH of a 10aM K+ sample. 4.3
Effect of the Carrier Complexation Strength Equations I, 2 and 3 indicate that the ion-exchange equilibrium is also influenced by the degree of complexation between the proton ionophore and the hydrogen ion. It is expected that an increase of the dissociation constant of the proton carrier would increase the potassium ion partitioning (Eq. 4) and decrease the proton transport process within the membrane phase (Eq. 5). This is proven by the experimental results of the CHEMFET. From Figure 7 it can be seen that the proton carrier dissociation constant (tridodecylamine with Ka 10 "1°"6 [50], ETH 2439 with Ka 10 -7.7 [51]) and ETH 1907 with Ka 10 "6"1 [52]) has a drastic effect on the linear range of the sensor response to protons. In addition, the potassium sensitivity of tridodecylamine-based CHEMFETs (36.5 mV/[K+]) has the lowest sensitivity when compared with the ETH 2439 based one (52.7 mV/[K+]) or better the ETH 1907 based sensor (55.1 mV/[K+]) 4.4
Effect of Ion Partitioning Another parameter influencing both the potassium ion partitioning and the proton transport process is the partition of the corresponding ions in the membrane phase. The partition of the hydrophilic ions from the aqueous sample
404
Ion-Partitioning M e m b r a n e s As Electroactive E l e m e n t s For T h e D e v e l o p m e n t O f A N o v e l Cation-Selective C H E M F E T Sensor S y s t e m Moschou
to the lipophilic membrane is expressed by the partition coefficient (k). The increase of the dielectric constant of the plasticizer used decreases the effective membrane lipophilicity. This will enhance the extraction of the hydrophilic ions 400 i¸ j
300
E 200
__ • i ~
100 !ii¸
0 "" ~ 2.5
0
,.--
o..~ IU.O ~ ~ -
pH
b~
"
Figure 7. Effect of the proton ionophore acidity (LogK~) on the potentiometric response of a K-CHEMFET varying the pH of a 104M K + aqueous sample.
in the corresponding membrane. The use of plasticizers of high dielectric constants are therefore expected to increase the amount of protons transporting into the ion-partitioning membrane. Additionally, the decrease of the dielectric constant of the membrane is expected to enhance the sensitivity of the corresponding CHEMFET to potassium ions, according to equation 2. The theoretically expected effect of the plasticizer dielectric constant on the potentiometric response of the K-CHEMFET is confirmed experimentally in Figure 8. The decrease of the plasticizer dielectric constant from o-NPOE (c: 23.9 [53]) to TBP (g: 7.95 [54]) and finally to DOS (g: 3.9 [53]) hinders the proton transport process decreasing the CHEMFET's potentiometric response to hydrogen ions. In addition, it increases the potassium sensitivity of the sensor to 55 mV/[K +] for the DOS-based sensor from 40 mV/[K +] or even negligible potentiometric response obtained from the TBP and o-NPOE-based sensors. The polymer matrix is another parameter affecting the physical and chemical properties of the membrane phase. It is expected that decreasing the ionic character of the membrane would result in a decrease of the ion partitioning and thus the proton transport into the polymeric membrane. In Figure 9 the proton response of the CHEMFETs with ion- partitioning membranes based on the polymer matrixes according to their electrical resistance, silicone rubber, PVC-NH2, PVC and PVC-COOH [55-57] is shown.
405
Ion-Partitioning Membranes As Electroactive Elements For The Development Of A Novel Cation-Selective CHEMFET Sensor System Moschou
From this data it can be seen that the increase of the electrical resistance of the membrane hinders the proton extraction to the bulk of the membrane phase resulting in the decrease of the CHEMFET's linear range of response to protons. In addition, the PVC-COOH is able to enhance the proton extraction in
250 200 ~'150 100DOS I'BP NPOE
50 0 2.5
4.0
5.5
7.0 pH
8.5
10.0
Figure 8. Effect of the plasticizer lipophilicity on the potentiometric response of the KCHEMFET varying the pH of a 104M K+ aqueous sample. the bulk of the membrane due to the protonated carboxylated groups [58] (which is not the case for the unprotonated groups of PVC-NH2 [59]) improving the CHEMFET's proton response. According to equation 2, any decrease in the H ÷ sensitivity will result in an increase of the potentiometric response of the sensor to the potassium ions. The performance of the K-CHEMFETs increasing the electrical resistance of the polymer matrix is attributed to the favored ion-exchange mechanism of K + with H + in the bulk of the membrane to the expense of the co-extraction mechanism of K ÷ with sample anions (CI). The co-extraction mechanism can also take place in the bulk of the membrane. Generally, it is not favored due to the limited partitioning of the C1- (absence of C1- carrier in the membrane). In the case of membranes based on polymer matrixes with low electrical resistance, the extraction of free chloride anions is more favorable, limiting the cation exchange mechanism of K + with H +. Therefore, the proton flux displaced from the bulk of the membrane to the proton transducer is limited decreasing the K-CHEMFET's potentiometric response to potassium ions (Figure 9).
406
Ion-Partitioning Membranes As Electroactive Elements For The Development Of A Novel Cation-Selective CHEMFET Sensor System Moschou
It is thus concluded that the potassium sensitivity of the K-CHEMFETs based on ion-partitioning membranes is increased by increasing the relative concentration of the analyte carrier relative to the proton counterpart in the polymeric membrane, the use of an acidic proton ionophore (high dissociation constant), the utilization of a low dielectric constant (lipophilic) plasticizer, and high electrical resistance polymer matrix. From these results the optimized potassium selective ion-partitioning membrane is developed using the proton
300 '~ 200 100
2..'
J.
©
Figure 9. Effect of the polymer matrix electrical resistance on the potentiometric response of the K-CHEMFET varying the pH in a 104M K + aqueous solution. carrier ETH 1907 and valinomycin at a molar ratio of 1:1, the plasticizer DOS, and silicone rubber as the polymer matrix. The K-CHEMFET based on the optimized ion-partitioning membrane responds to the increase of the potassium ion activity in the sample with a fast response time (in the order of seconds) and the theoretical sensitivity of 59 mV/decade [K+] [42]. Based on the findings of the studies on the development of the potassium selective CHEMFET, we proceeded in the development of a CHEMFET sensitive to the multivalent calcium ions. The calcium-selective CHEMFET was based on an ion-partitioning membrane doped with the proton carrier ETH 1907 in equimolar concentration with the calcium ionophore ETH 1001. The polymer matrix and the plasticizer combination used was the PVC (of low electrical resistance) with o-NPOE (of high dielectric constant) that have been successfully used for the construction of calcium selective polymeric membranes due to the enhanced calcium ion partition they offer [1,2]. Even
407
Ion-Partitioning Membranes As Electroactive Elements For The Development Of A Novel Cation-Selective CHEMFET Sensor System Moschou
though anyone would expect that the sensor based on the ion-partitioning membrane with the above composition should work, this was not the real case. The Ca-CHEMFET with this membrane showed a transient response to the increase of the calcium ion activity in the sample. The magnitude of this transient response was super-Nemstian, in the order of 39mV/decade [Ca2+], which verifies the direct measurement of the membrane proton flow by the pHsensitive transducer. But the transient response indicates a short-term response contradicting a long-term response limiting the desired the membrane proton flow towards the pH-sensitive transducer. In order to assist the ion-exchange mechanism of calcium ions with protons we dramatically increased the Ca 2+ to H + ionophore molar ratio in the membrane up to a value of 15. This effect increased the magnitude of the transient response but did not stabilize the CaCHEMFET's signal. Based on these results it was postulated that there must exist another mechanism that is in conflict with the desirable Ca 2+ - H + ion-exchange. Such a mechanism could be the cation/anion coextraction. In such a case, the membrane proton flow to the pH-ISFET's gate is eliminated,
Ca2+(aq)+2L(org) +2ax
~--~CaL22+(org)+
2X'(org)
(6)
where CaL22+, L are the complexed and free calcium ionophore. According to equation 6 the partition of the counteranions is free (not mediated by an ionophore) and is therefore totally dependent on the polymeric membrane characteristics. The higher the dielectric constant of the plasticizer and the lower the electrical resistance of the polymer matrix the more favorable the coextraction mechanism becomes to the expense of the ion-exchange mechanism. For these reasons PVC/o-NPOE -based membrane was eliminated in favor of silicone rubber, which is known for its high electrical resistance while the use of an extra plasticizer as the membrane organic solvent is not required. The above change in the membrane chemical composition had a dramatic result on the CHEMFET's potentiometric response. The Ca-CHEMFET based on the new ion-partitioning membrane presented a stable response to the increase of the calcium ion activity in the sample with a fast response time of less than a minute and the expected super-Nernstian sensitivity of 55 mV/decade [Ca2+]. In order to verify that the observed changes were actually due to the counteranion extraction in the membrane the potentiometric response of the CaCHEMFET was evaluated in solutions with different counterions. Increasing the lipophilicity of the counterion would result in an increase in the coextraction efficiency, and thus a decrease in the proton flux. This potentiometric response of the Ca-CHEMFET is presented in Figure 10. The increase of the counteranion lipophilicity (equation 2) increases the unmediated counteranion partition in the
408
Ion-Partitioning Membranes As Electroactive Elements For The Development Of A Novel Cation-Selective CHEMFET Sensor System Moschou
membrane therefore favoring the co-extraction mechanism. So, less amount of calcium ions are ion-exchanged with protons. The result is the decreasing membrane proton flow towards the pH-ISFET diminishing the CHEMFET's potentiometric response. The effect of the counteranion co-extraction is also expected to be less pronounced after the membrane preconditioning with a low pH sample solution. According to equation 3, the membrane proton loading will favor the ion-exchange mechanism by shifting the ion-exchange equilibrium towards the calcium ion
30
;;* 20
10
CI"
NO 3-
I-
SCN" C10 4"
CI"
Cl-prot
Figure 10. Effect of the counteranion lipophilicity on the potentiometric response of the Ca-CHEMFET in a pH 7.0 sample solution containing a 102M concentration of the background electrolyte. [Cl'prot denotes to the study of the CHEMFET's response in the presence of chloride counteranions in the sample after the conditioning of the sensor in a pH 3.0 preconditioning solution]. partition. The experimental verification of this effect is presented in Figure 10 where the memory effect of the membrane after the partition of the lipophilic perchlorate anions is erased by the membrane preconditioning in a KHP solution at pH 3.0. The Ca-selective CHEMFET based on an ion-partitioning membrane has been furthermore optimized by the use of a thin dextran layer between the pHISFET's gate and the ion-partitioning membrane. The role of this layer is to stabilize the thin aqueous layer that is formed by the diffusion of water molecules through the polymeric membrane. The optimized Ca-CHEMFET is based on an ion-partitioning membrane of silicon rubber with a molar ratio of
409
Ion-Partitioning Membranes As Electroactive Elements For The Development Of A Novel Cation-Selective CHEMFET Sensor System Moschou
ETH 1001 to ETH 1907 equal to 15 and which incorporates the lipophilic anionic sites potassium tetrakis [3,5-bis(trifluoromethyl) phenyl] borate. The evaluation of the analytical performance of the potassium and calcium-selective CHEMFETs based on the optimized ion-partitioning membranes has been carried out by the potassium and calcium ion determination in blood serum samples. The K + and Ca z+ activity of the samples has been also determined by the use of the K + and Ca2+-selective ISEs as the reference method. In Figure 11 the correlation graphs of the determination of the monovalent potassium and divalent calcium ions is presented with a slope of 1.05 for the potassium and 0.936 for the calcium determination indicating a good agreement between the results obtained by the CHEMFETs and the reference method. 12.0 10.0 [.-,
z~
8.0 6.0
r,.)
4.o
2.0 2.0
4.0
6.0
8.0
10.0
12.0
z+
[I ]ISEX 104M Figure 11. Correlation graph of the potassium (e) and calcium (1"1)ion determination in blood serum samples by the CHEMFETs based on ion-partitioning membranes using the K+ and Ca2+ -selectiveISEs as the reference method. 5.
CONCLUSIONS
The development of a novel type of bulk active membranes as electroactive elements in CHEMFET design is shown. The potentiometric sensor system based on these membranes follows a new response mechanism, which has been fully demonstrated. The sensor is based on a bulk active ionpartitioning membrane as the chemical recognition element while the signal transduction is carried out by a pH-ISFET. Due to the chemical composition of the ion-partitioning membrane the ion-exchange of analyte cations with protons
410
Ion-Partitioning Membranes As Electroactive Elements For The Development Of A Novel Cation-Selective CHEMFET Sensor System Moschou
takes place in the bulk of the membrane. The increase of the analyte activity in the sample results in the analyte cation partition in the bulk of the membrane. The electroneutrality principle is fulfilled throughout the bulk of the membrane due to the simultaneous displacement of protons of equal charge out of the membrane. The outer membrane environment where the protons are expelled is both the aqueous sample, in the vicinity of the outer phase boundary of the membrane, and the thin aqueous layer on the pH-ISFET's gate, that lies next to the inner phase boundary of the membrane. Therefore, the direct measurement of the proton flow expelled by the membrane is the indirect determination of the analyte ion activity in the sample. The effect of the chemical constituents of the polymeric membrane and the chemical composition of the aqueous sample on the sensor response has been evaluated. It has been shown that the increase of the analyte with respect to the proton ionophore content in the membrane and the use of an acidic proton ionophore increase the membrane proton flow towards the pH-ISFET's gate. In addition, the use of a plasticizer of low dielectric constant and a polymer matrix of high electrical resistance favor the ion-exchange mechanism of analyte cations with protons stabilizing the CHEMFET's signal. Finally, the increase of the analyte cation activity, the increase of the sample pH and the use of counteranions of low lipophilicity in the sample result in the increase of the potentiometric response of the sensor towards the analyte cations. The great advantage of the CHEMFET based on the presented response mechanism is the fact that the sensitivity of the sensor is not affected by the charge of the analyte ion but is rather constant and equal to 59.2 mV/decade of the cation. This is because the potentiometric response of the CHEMFET is actually developed by the measurement of the protons displaced out of the ionpartitioning membrane. Therefore, the actual sensitivity of the sensor is 59.2mV/pH for every cation analyte determined, and is translated into the sensitivity of 59.2mV/[IZ+]. Therefore, the potentiometric sensor based on the demonstrated original response mechanism presents a sensitivity that is not proportional to the inverse of the analyte cation charge, as for the rest of the potentiometric sensors presented so far in the literature. The importance of the developed sensor is more pronounced in applications where the multivalent cation determination in the presence of high concentration of monovalent interfering cations is desired. 6.
ACKNOWLEDGMENTS
We gratefully acknowledge Thermo Orion for providing the pH-ISFETs, the pH-electrode and the pulseless pump. This work was supported by the Greek General Secretariat of Research. 411
Ion-Partitioning Membranes As Electroactive Elements For The Development Of A Novel Cation-Selective CHEMFET Sensor System Moschou
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INDEX Arsenic 344 2-D fourier-transform 90 2-D phase-contrast 90 Adsorption-based contactors 158 Affinity membranes 246 Artificial kidney 220 Bio-artificial hybrid systems 187 Bioartificial organs 197 Biocatalytic membrane reactor 241 Biocatalytic membrane reactors 188 Biocompatibility of sensors 385 Biofunctional membranes 233 Biopharmaceuticals 283 Bioreactor coupled to membrane processes 244 Biosensors 379 Biotechnological and medical applications 187 Cake filteration 28 Cation-selective sensor 393 CHEMFET 393 Chitin binding domain 265 CO2 removal 131 CO2-production for greenhouses 132 Contactors as reactors 158 Cooling curves 53 Countercurrent flow 91 Cytocompatibility 207 DSC 48 Electrochemical polymerization 304 Electron beam 308 Electron paramagnetic resonance 237 Emulsion enzyme membrane reactor 194 Emulsion pertraction 140 Flux decline 28 Fouling 27, 65 Functionalized membranes 329 Gel-filled nanofiltration membranes 353 Gradient-echo pulse sequence 95 Graetz number 16 Hemodialysis membranes 219 Hemodialyzers 112 High-speed membrane adsorbers 283 Hollow fiber contained liquid membrane 152 Hollow fiber membrane modules 5 Immobilized enzyme 189 In-situ crosslink 309 Interfacial tension 150 Ion-partitioning membrane 398 Ionophore 395 ),-irradiation 306 Liquid-gas contactors 153 Liquid-liquid contactors 156 Membrane adsorber modules 285 Membrane aromatic recovery systems 166 Membrane biocompatibility 209 Membrane contactors 125, 147 Membrane gas absorption 126
Membrane immobilization 233 Membrane stripping 142 Membranes for the immobilization of biomolecules 380 Module configurations 151 MRI-based flow imaging 90 Nanofiltration 353 Pertraction 136 Plasma 308 Polymer grafted membranes 299 Pore blockage - cake filtration model 29 Pore structure 30 Responsive membranes 311 Rheology 48 Selectivity control through membranes 382 Sherwood number 16 Silanization and polymerization 309 Site-specifically immobilized enzymes 237 Spin-echo pluse sequence 92 Spin label titration 238 Supercritical fluid-based contactors 157 Surface active membrane 394 Thermally-induced phase-separation 45 Tunable separations 329 Ultrasonic time-domain reflectometry 66 UV-initiation 306 Viscometry 56 Whole polymer grafting 310
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