Faculteit Bio-ingenieurswetenschappen
Academiejaar 2013 - 2014
Performance and fouling behaviour of novel generations of osmotic membranes in forward osmosis
Niels Lefebure Promotor: Prof. dr. ir. Arne R.D. Verliefde Tutor: Machawe M. Motsa
Masterproef voorgedragen tot het behalen van de graad van Master in de bio-ingenieurswetenschappen: Milieutechnologie
i
Preface
When I started last summer, I could have never imagined how bumpy the road would be that was lying in
front of me. Never before I’ve had experienced so many altering moments of joy and despair
consecutively. Trying to be a part of scientific research has proven to be very challenging, educational,
often tiresome, but in the end very satisfying. Along the way, many people have been of great importance
to the eventual success of this thesis. In order to entirely complete this work, it is therefore time to thank
them.
I’ll start with the man who made all of this possible, Arne. I can’t imagine a professor who is closer to his
students than him. He was accessible, supportive, and gave the necessary freedom to establish my own
ideas. Furthermore, he gave me the opportunity to perform a big part of my thesis in Johannesburg,
South Africa, which has been a life changing experience. It broadened the purpose of the thesis to more
than just scientific work, which I appreciate very hard. As if this intercultural experience wasn’t enough, I
had the pleasure to be guided throughout the year by a magnificent man from Swaziland, Machawe. In
moments of despair, he was the one who made me come back to lab day after day. Together with Oranso,
they prepared me for the South African experience. Once on the other side of the world, Sabelo took over
from Arne by providing us with the necessary support. Although even he couldn’t change the South
African rhythm, which has been testing my patience thoroughly, he did whatever he could to help us.
Next to him, Nsika, Gcina, Sako, Richard, Patrick, and many others have made me feel at home in that
beautiful country. I really hope I will be able to see them once again. To make the experience complete I
was accompanied by Ben, Kwinten, and Wouter. By supporting each other, having a lot of fun together,
and being able to get things off our chest once in a while, we were a great team (go Fantastic 4!).
Also in Ghent, a whole bunch of great people have been essential. Seba helped me with the start-up (for
which I just realise I still owe you a beer!). Arnout, Marjolein, Klaas, Eric, and Quenten were also
accessible all the time. They made me realize that completing a thesis can be done by combining hard
work and having a lot of fun. Next to them, many others also contributed to the great atmosphere in the
lab, which is definitely one of the strengths of the rather small and cozy department PaInT.
Of course I can’t forget to thank my parents, which have supported me throughout the whole year. As
compassionate as they are, I can image what they have been through when I was having a hard time. I
can’t thank them enough to finance and support me when I was in Jozi for nearly a full semester. Without
my brother Siebe, I would not have been able to relativise when I needed to (dank u broereman!).
It always seem impossible until it’s done – Nelson Rolihlahla Mandela.
Niels
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Abstract
Water scarcity is currently one of society’s most persistent problems, affecting over 1.2 billion people
worldwide. One way of providing access to potable water is accomplished by seawater desalination,
currently dominated by the reverse osmosis (RO) process. This technique has led to a drastic
reduction of the overall energy demand of desalination processes. However, 2 kWh/m³ is still
required for the reverse osmosis system, covering nearly half of the overall costs of the entire plant.
One possibility to tackle this energy issue is offered by an energy-efficient, stand-alone forward
osmosis (FO) desalination system (although under certain conditions). Nevertheless, the economic
viability of the stand-alone process is strongly dependent on the overall yield of the system, which is
in turn related to the permeate flux reached by the semi-permeable membrane. Since the extent of
water permeation is mainly governed by the performance parameters and fouling propensity of the
applied membrane, both traditional (CTA) and novel (TFC and POR) membranes were examined
under (combined) fouling conditions throughout this study. The first aim was to establish a
fundamental understanding of the mechanisms governing organic and colloidal fouling (both
individually and combined) at the high ionic strength of seawater, using sodium alginate and silica
colloids as model foulants. Consequently, a comparison was made between the state-of-the-art
membranes in order to elucidate the importance of foulant-membrane interaction as well as to
assess their applicability in reality. In all cases, the extended Derjaguin-Landau-Verwey-Overbeek
(XDLVO) surface energy analysis was applied to further investigate fouling mechanisms and to
examine the link between flux decline rates and foulant-foulant and foulant-membrane interactions.
Our results indicate that substantial losses in permeate flux occurred in both the individual (alginate
and silica colloids separately) and combined cases when the ionic strength of the feed solution was
elevated to the level of seawater (0.5 M, mimicked by 0.476 M NaCl and 0.008 M CaCl2), mainly
caused by the presence of calcium (Ca2+). Even though the total concentration of the foulants was 1.2
g/L (1g/L silica colloids and 0.2 g/L alginate) when combined, a similar flux decline trend to that of
individual alginate fouling could be observed, suggesting that the alginate-calcium complexation
possibly overwhelmed the presence of silica colloids. Further research on this matter was performed
by applying different sets of (newly proposed) consecutive fouling experiments, confirming to a
certain extent that the alginate-calcium complextation was indeed dominant during combined
fouling. Next, the POR membrane was shown to outperform the CTA and TFC membranes in terms of
water permeability and permeate flux, while being far less resilient to fouling with either alginate,
silica colloids, and their combination compared to its counterparts (CTA and TFC). By assessing the
overall membrane performance relatively to each other, it was concluded that currently, none of the
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tested membranes was fully suitable for the stand-alone seawater desalination process. Expanding
on the extended Derjaguin-Landau-Verwery-Overbeek (XDLVO) approach, it was shown that the
calculated adhesive energies largely matched with the observed experimental dissimilarities (in
terms of losses in permeate flux decline) for each foulant-membrane combination during the initial
stages of the fouling runs. Once a fouling layer was formed, the cohesive forces were shown to be of
major importance for further membrane fouling. Based on the XDLVO calculations, the relatively high
fouling propensity of POR could be mainly attributed to its surface functionality.
v
Table of contents
Preface ......................................................................................................................................... i
Abstract ..................................................................................................................................... iii
Table of contents ........................................................................................................................ v
List of figures ............................................................................................................................. ix
List of tables ............................................................................................................................. xiii
List of abbreviations ................................................................................................................. xv
List of symbols ......................................................................................................................... xvii
Part I: Introduction .................................................................................................................... 1
1. Problem statement .......................................................................................................... 1
2. Goal of the study ............................................................................................................. 1
3. General Outline of the thesis .......................................................................................... 2
Part II: Literature study .............................................................................................................. 3
Chapter 1 – osmotically driven membrane processes ........................................................... 3
1.1. Forward osmosis ...................................................................................................... 4
1.2. Pressure retarded osmosis ....................................................................................... 6
Chapter 2 - Factors influencing the performance of osmotically driven membrane
processes................................................................................................................................. 9
2.1. Concentration polarisation ...................................................................................... 9
2.2. Reverse solute diffusion ......................................................................................... 12
2.3. Membrane properties ............................................................................................ 12
2.4. Membrane orientation ........................................................................................... 14
2.5. Operating conditions .............................................................................................. 15
2.6. Fouling .................................................................................................................... 17
Chapter 3 - Membrane development in ODMPs .................................................................. 23
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3.1. Types of membranes ............................................................................................. 23
3.2. Fabrication methods .............................................................................................. 24
3.3. Membrane design and development .................................................................... 25
3.4. Characterisation of fouling (propensity) ............................................................... 27
Part III: Materials and methods .............................................................................................. 29
1. Experimental determination of membrane performance in ODMPs ........................... 29
1.1. Membranes ............................................................................................................ 29
1.2. Foulants ................................................................................................................. 29
1.3. Feed solution chemistry ........................................................................................ 30
1.4. Experimental set-up and protocol ......................................................................... 30
1.5. Determination of water flux – Performance evaluation ....................................... 32
2. Characterisation of membranes and foulants .............................................................. 33
2.1. Surface tension measurements using contact angles ........................................... 33
Part IV: Results and discussion ............................................................................................... 35
1. CHaracterisation of Forward Osmosis membranes ...................................................... 35
1.1. Determination of the membrane performance parameters ................................ 35
1.2. Determination of the membrane hydrophilicity ................................................... 35
2. Determination of mechanisms governing combined organic and colloidal fouling in
Forward osmosis .................................................................................................................. 37
2.1. Single and combined membrane fouling ............................................................... 37
2.2. Consecutive fouling ............................................................................................... 41
2.3. Foulant-foulant and foulant-membrane interaction energies .............................. 45
3. Comparison of performance and fouling behaviour of traditional and novel forward
osmosis membranes ............................................................................................................. 49
3.1. Single and combined fouling ................................................................................. 49
3.2. Relation between interfacial free energies of adhesion and membrane fouling . 52
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3.3. Membrane applicability in stand-alone FO desalination ....................................... 53
Part V: Conclusions and future prospects ............................................................................... 55
Bibliography .............................................................................................................................. 59
Appendices ............................................................................................................................... 67
Appendix A. Contact angle measurements .......................................................................... 67
Appendix B. Determination of A, B, R, and S ........................................................................ 69
Appendix C. SEM images of clean membranes .................................................................... 71
Appendix D. (SEM) Images of fouled membranes ................................................................ 73
Appendix D. Determinition of zeta potential and hydrodynamic particle size .................... 76
D1. Particle sizing .............................................................................................................. 76
D2. Zeta potential measurements .................................................................................... 76
D3. Results ......................................................................................................................... 76
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ix
List of figures
Figure 1. Schematic representation of osmotically driven and pressure driven membrane processes.
Respectively from left to right: forward osmosis (FO), assisted forward osmosis (AFO), pressure
retarded osmosis (PRO), and reverse osmosis (RO). The low saline solution is displayed in light blue
(this is pure water in case of RO), while the high saline solution is shown in dark blue. This figure was
adapted from Cath, et al. [9]. ...................................................................................................................3
Figure 2. Schematic representation of the application of FO in seawater desalination by (A) a stand-
alone process using an artificial draw solution in a closed-loop system or (B) a hybrid with RO, in
which the seawater (brine) is diluted by the FO process. The symbol “O” represents a low pressure
pump, which is required for the circulation. These figures were adapted from McCutcheon, et al. [16]
and Bamaga, et al. [14] respectively. .......................................................................................................5
Figure 3. Schematic diagram of osmotic power generation. As the water permeates from the feed to
the draw solution side, the latter is diluted. One part of this still pressurised draw solution then
passes a hydropower turbine while the other part is sent through a pressure exchanger to pressurise
the incoming draw stream (seawater). The figure was adapted from Achilli, et al. [23]. .......................7
Figure 4. Water flux (Jw) and power density (W) versus applied hydraulic pressure (ΔP) comparing
ODMPs with RO. Both Jw and ΔP are orientated from the draw to the feed solution. This graph was
adapted from Achilli, et al. [23]. ..............................................................................................................8
Figure 5. Illustration of osmotic driving force profiles for osmosis for (a) a dense symmetric
membrane, (b) an asymmetric membrane with the AL-DS configuration (PRO mode), and (c) an
asymmetric membrane AL-FS configuration (FO mode) [34]. .............................................................. 11
Figure 6 Schematic representation of the coupled influence of hydrodynamic forces (y-axis) and
intermolecular adhesion (x-axis) on membrane fouling in the presence of alginate, AHA and BSA
respectively. The first row illustrates the effects in AL-FS mode, while the lower two represent the
situation in AL-DS mode [53]. ............................................................................................................... 18
Figure 7. SEM images of the cross-section of TFC (PRO) hollow fiber [83] and flat sheet (FO)
membranes [87], respectively left and right. ........................................................................................ 24
Figure 8. Schematic representation of the phase inversion process by immersion precipitation. .. Fout!
Bladwijzer niet gedefinieerd.
Figure 9. Schematic representation of the Interfacial polymerization processFout! Bladwijzer niet
gedefinieerd.
Figure 10. Schematic representation of the bench-scale experimental set-up which mimics a stand-
alone FO desalination system.. ............................................................................................................. 31
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Figure 11. Permeate flux decline curves in case of fouling a CTA membrane with single foulants (1 g/L
silica colloids (ST-ZL) and 200 mg/L alginate (ALG)) and their combination (Comb) in the absence of
mono and divalent salts. The draw solution was adjusted to give a similar initial flux to that of fouling
in the presence of background electrolytes. Each fouling experiments was performed in AL-DS mode
at a cross-flow of 30 l/h for 24 h. .......................................................................................................... 39
Figure 12. Permeate flux decline curves in case of fouling a CTA membrane with single foulants (1 g/L
silica colloids and 200 mg/L alginate) and their combination in the presence of mono and divalent
salts (0.476 M NaCl and 0.008 M CaCl2). Each fouling experiment was performed in AL-DS mode at a
cross-flow of 30 l/h for 24 h. ................................................................................................................. 40
Figure 13. Permeate flux decline curves in case of combined fouling a CTA membrane (200 mg/L
alginate and 1 g/L silica colloids) with and without the 0.008 M CaCl2. Each fouling experiment was
performed in AL-DS mode at a cross-flow of 30 l/h for 10 h. ............................................................... 40
Figure 14. Flux decline curves in case of layer by layer membrane fouling of CTA membranes with
single foulants– in alternating sequences – and a combination of both foulants. A fresh feed solution
with a total ionic strength of 0.5 M was used for the subsequent fouling run. Each fouling experiment
was performed in AL-DS mode at a cross-flow of 30 l/h for 24 h. ........................................................ 42
Figure 15. Flux decline curves in case of layer by layer membrane fouling of CTA membranes with
single foulants – in alternating sequences – and a combination of both foulants. For the subsequent
fouling run, the total ionic strength of the feed solution was adjusted to match the total ionic
strength in the end of the previous run. Each fouling experiment was performed in AL-DS mode at a
cross-flow of 30 l/h for 24 h. ................................................................................................................. 43
Figure 16. Flux decline curves in case of consecutive membrane fouling of CTA membranes with silica
colloids. For the subsequent fouling run, the total ionic strength of the feed solution was adjusted to
match the total ionic strength in the end of the previous run. Each fouling experiment was performed
in AL-DS mode at a cross-flow of 30 l/h for 24 h. .................................................................................. 44
Figure 17. Overview permeate flux decline curves in case of fouling a (a) CTA (AL-FS mode), (b) TFC,
(c) CTA (AL-DS), and (d) POR membrane with single foulants (1 g/L ST-ZL silica colloids and 200 mg/L
alginate) and their combination in the presence of mono and divalent salts (0.476 M NaCl and 0.008
M CaCl2). Each fouling experiment was performed in AL-FS mode (except the CTA membrane
depicted in (c)) at a cross-flow of 30 l/h for 24 h. ................................................................................. 51
Figure 18. Picture of the Krüss DSA 10-Mk2 contact angle measurement device with a camera (1),
sample platform (+ sample) (2), needle (for droplet deposition) (3), and light source (4). .................. 67
Figure 19. Illustration of the contact angle calculation principle. The angle is measured between the
baseline (blue) and the tangent through the point where the droplet touches the surface (red). ...... 68
xi
Figure 20. SEM images from the active layer (a), porous layer(b), and cross section (c) of a CTA
membrane. ............................................................................................................................................ 71
Figure 21. SEM images from the active layer (a), porous layer(b), and cross section (c) of a TFC
membrane. ............................................................................................................................................ 72
Figure 22. Images of fouled (a) POR, (b) CTA, and (c) TFC membranes (in AL-FS mode) when both
alginate and silica colloids were applied in the highly saline feed solution. ........................................ 73
Figure 23. Images of samples prepared for contact angle analysis of combined fouling with alginate
and silica colloids. .................................................................................................................................. 73
Figure 24. Top view SEM images of CTA membrane fouled with alginate (a), silica colloids (b), and
their combination (c). In all cases, the feed solutions contained 0.476 M NaCl and 0.008 M CaCl2. ... 74
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xiii
List of tables
Table 1. Overview membrane properties and performance parameters from recent developments in
FO membranes. ................................................................................ Fout! Bladwijzer niet gedefinieerd.
Table 2. Overview different sets of experiments fouling experiments performed on each type of
membrane. In case of alginate, experiments were performed in both the FO and PRO mode. .......... 31
Table 3. Overview of the membrane performance parameters and hydrophilicity for each membrane
type. ...................................................................................................................................................... 36
Table 4. Initial flux decline rates for each run during the sequential fouling experiments. ................. 42
Table 5. Contact angles ( of H2O + 0.476 M NaCl + 0.008M CaCl2) and surface free energies for the
single foulants and their combination. ................................................................................................. 47
Table 6. angles (H2O + 0.476 M NaCl + 0.008M CaCl2) and surface free energies for the different
membranes. .......................................................................................................................................... 47
Table 7. Foulant-foulant interfacial free energies of cohesion and foulant-membrane/foulant-foulant
interfacial free energies of adhesion. CTA-PRO refers to the case were the CTA membrane was
applied in AL-DS mode, while CTA-FO indicates the membrane was applied in AL-FS mode. ............. 47
Table 8. Respective draw concentrations and final flux decline percentages (compared to the baseline
after 10 h) for each membrane (in AL-FS mode). ................................................................................. 50
Table 9. Overview initial flux decline rates for each foulant-membrane combination. ....................... 51
Table 10. Assessment of membrane performance (in AL-FS mode) based on two criteria: (1) intrinsic
parameters (A, B, S, and R) and (2) fouling propensity. (1) is rated as bad, average or good, while (2) is
rated as low, medium or high. .............................................................................................................. 54
Table 11. Measured zeta potential and hydrodynamic diameter of alginate and silica colloids in
different solutions. ................................................................................................................................ 76
xiv
List of abbreviations
AB Lewis acid-base
AFM Atomic force measurement
AFO Assisted forward osmosis
AL-DS Active layer facing the draw solution
AL-FS Active layer facing the feed solution
ALG (Sodium) Alginate
BSA Bovine serum albumin
CECP Concentrative external concentration polarisation
CICP Concentrative internal concentration polarisation
COMB Combined
CP Concentration polarisation
CTA Cellulose triacetate
CTA-W Woven cellulose triacetate
DECP Dilutive external concentration polarisation
DI De-ionised water
DICP Dilutive internal concentration polarisation
DS(s) Draw solution(s)
ECP External concentration polarisation
EPS Extracellular polysaccharides
FO Forward osmosis
FS Feed solution
HA Humic acid
HTI Hydration Technologies Inc.
ICP Internal concentration polarisation
L-DOPA Poly amino acid 3-(3,4-Dihydroxyphenyl)-L-alanine
LMH Liter per square meter per hour
LW Liftshitz-van der Waals
NF Nanofiltration
ODMP(s) Osmotically driven membrane process(es)
PA Polyamide
PES Polyethersulfone
PET Polyethylene terephthalate
xvi
POR Membrane (confidential)
PRO Pressure retarded osmosis
PSf Polysulfone
RO Reverse osmosis
RSD Reverse solute diffusion
SEM Scanning electron microscopy
ST-ZL Silica colloids
TFC Thin film composite
UF Ultrafiltration
XDLVO Extended Derjaguin-Landau-Verwery-Overbeek
xvii
List of symbols
A Water permeability coefficient (L/m²/h/bar)
Membrane surface (m²)
B Solute permeability coefficient (m/s)
Draw solution concentration (moles/L)
Feed solution concentration (moles/L)
Permeate concentration (moles/L)
D Diffusion coefficient (m²/s)
Hydraulic diameter (m)
Diffusion coefficient of the draw solute (m²/s)
i Van't Hoff factor (-)
Js Solute flux (moles/m²/h)
Water flux (L/m²/h)
K Solute resistivity for diffusion within the porous support layer (-)
k Mass transfer coefficient of the bulk solution (-)
L Length (m)
M Molarity (moles/L)
m Mass (kg)
n Number of moles (moles)
R Rejection (%)
R Gas constant (L.atm/K/moles)
Re Reynolds number (-)
S Structure factor (µm)
Sc Schmidt number (-)
Sh Sherwood number (-)
t Time (s)
T Temperature (K)
V Volume (L)
Volume water permeated (L)
W Power desnity (W)
Maximum power density (W)
x Membrane thickness (m)
Electron-acceptor component
xviii
Electron-donor component
Liftshitz-van der Waals free energy component
Interfacial free energy of adhesion (acid-base component)
Interfacial free energy of adhesion (Lifshitz-van der Waals component)
Total interfacial free energy of cohesion
Total interfacial free energy of adhesion
Hydraulic pressure differential (bar)
Osmotic pressure differential (bar)
θ Contact angle (°)
Bulk osmotic pressure of the draw solution (bar)
Osmotic pressure of the draw solution at the membrane surface (bar)
Bulk osmotic pressure of the feed solution (bar)
Osmotic pressure of the feed solution at the membrane surface (bar)
Water density (kg/m³)
τ Tortuosity (-)
Porosity (-)
1
Part I: Introduction
1. PROBLEM STATEMENT Being the foundation of life, water is inevitably of paramount importance for humankind.
Nevertheless, over 1.2 billion people worldwide still lack access to clean and safe drinking water,
drastically limiting their standard of living [1]. As 97.5% of the world’s water capacity consists of salt
water [1], the amount of seawater desalination plants throughout the world has been continuously
rising during the last 50 years in order to counteract water scarcity [2]. Currently, reverse osmosis is
the most widespread technology amongst the established desalination processes [3]. By using a
semi-permeable membrane as a selective barrier for solutes and applying an external pressure to
make the water pass through it, a highly purified permeate can be obtained. Although reverse
osmosis is less energy intensive compared to thermal desalination processes such as multi-flash
distillation (~20 kWh/m³) [4], the energy demand is still considerably high (~2 kWh/m³) [5]. A possible
solution to tackle the latter is using osmotically driven membrane processes (ODMPs), which provide
an energy-efficient way of desalination or can be implemented in the reverse osmosis plant to
reduce the overall energy consumption (Figure 2). However, for these emerging technologies to
become economically viable, some major challenges need to be overcome. Two of these challenges
are (specific) membrane development and membrane fouling, both inextricably linked to the overall
performance of osmotically driven membrane processes. Even though membrane fouling is
considered to be lower in forward osmosis compared to reverse osmosis, it is still inevitable [6]. As
the extent of deposition and adhesion of foulants on the membrane largely depends on the
interaction between both, proper understanding of these interactions is imperative in membrane
design and development aiming to limit the adverse effect of membrane fouling on the overall
performance of ODMPs. Nevertheless, most of the proposed novel membranes for ODMPs have not
been examined under fouling conditions. In addition, the interactions between multiple foulants, as
occurring in real conditions, and the effect thereof on fouling, have not been investigated properly.
2. GOAL OF THE STUDY This study is aimed to establish a fundamental understanding of the mechanisms governing organic
and colloidal fouling of FO membranes. Sodium alginate and silica colloids were selected as
respective model foulants to perform single, combined, and consecutive fouling experiments on a
lab-scale. Using multiple foulants at once, which is a shortcoming in previous studies regarding
ODMPs, will provide insight in foulant-foulant interactions and their influence on fouling in FO.
Examining the nature of these foulant-foulant interactions more fundamentally is achieved through
2
(newly proposed) layer by layer membrane fouling performed during the consecutive fouling
experiments. In this study, both traditional (CTA) and novel (TFC and POR) membranes will be
examined. The former will establish a reference in terms of fouling behaviour to enable a proper
assessment of these novel generations of membranes, which is currently entirely lacking. Although
the state-of-the-art commercial membranes have resulted in a renewed interest in ODMPs, an
increase in membrane performance is imperative for those processes to eventually break through
completely. For this matter, all of the recent publications regarding membrane development have
been focussing on improving the permeability (to increase the product yield), while maintaining a
decent salt rejection (to ensure a good product quality). However, these membranes need to be
evaluated under fouling conditions as well to mimic real-life situations, which will be done in this
thesis for the first time. Moreover, using different membrane surfaces will provide insight in foulant-
membrane interactions. In all cases, the goal was to investigate fouling behaviour in seawater
desalination environments. Therefore, examination of the fouling behaviour was performed at the
high ionic strength of seawater, which enables the assessment of the applicability of ODMPs in reality
(i.e. the stand-alone FO desalination system). In the last part of the study, characterisation of
membrane fouling using the extended Derjaguin-Landau-Verwey-Overbeek (XDLVO) surface energy
analysis to obtain more fundamental data regarding the foulant-membrane and foulant-foulant
interactions. In the end, this thesis is meant to acquire more fundamental insight in membrane
fouling, to evaluate novel generations of membranes, to relate performance to membrane surface
and foulant characteristics as determined using novel methodologies (i.e. consecutive fouling), and
ultimately to support the development of the promising ODMPs, by outlying which membrane and
foulant characteristics are predominant.
3. GENERAL OUTLINE OF THE THESIS This work is divided into five parts of which the first part is this introduction. The second part (i.e. the
literature study) is subdivided into 3 chapters, gradually going into more detail. The initial chapter in
the literature study introduces the principles and applications of the different osmotically driven
membrane processes. A second chapter will reveal the key factors that influence the performance
and viability of these emerging technologies. Due to the relevance for this study, membrane fouling
will be the main focus of this chapter. The third and final chapter of the literature study is meant to
briefly review the research related to membrane development in osmotically driven membrane
processes. The third part of the thesis consists of the materials and methods employed to investigate
the abovementioned objectives. Subsequently, the experimental results and data regarding the
characterisation are discussed and interrelated in the fourth part. Finally, the fifth part outlines the
general conclusions of the thesis as well as recommendations for future research.
3
Part II: Literature study
CHAPTER 1 – OSMOTICALLY DRIVEN MEMBRANE PROCESSES Reverse osmosis (RO) is the world’s most widespread type of membrane-based seawater
desalination processes. It relies on a semi-permeable membrane, which allows the solvent to pass
while the solutes are rejected. These RO membranes usually reject up to 98-99.5% of the solutes,
resulting in a very pure permeate and a concentrate with a high salt content. However, an external
force is required to make the water permeate through the polymeric structure of the membrane.
This is achieved by applying a hydraulic pressure on the feed stream, which is mainly seawater or
brackish water in case of RO (Figure 1). In seawater desalination, pressures ranging from 55 to 80 bar
are common, depending on the osmotic pressure of the seawater [2]. It is self-evident that applying
such high pressures implicates a high energy demand, making RO desalination a very energy
intensive process. Although a drastic reduction in energy consumption has been achieved - for the
moment, 2.2 kWh/m³ is typically required for seawater RO at 50% recovery, compared to 20 kWh/m³
in the seventies [7] - the energy demand still accounts for 40-50% of the total costs of the RO
desalination plant [8]. In addition, the disposal of the concentrated brine can have detrimental
effects on the local marine life and water quality due to the high salt content and the presence of
chemicals for cleaning purposes. For such reasons, research has been looking for more energy-
efficient and less environmental disturbing alternatives to RO, leading to the re-discovery of
osmotically driven membrane processes (ODMPs). These can be subdivided into forward osmosis
(FO), pressure retarded osmosis (PRO), and assisted forward osmosis (AFO), all of which are briefly
discussed below.
Figure 1. Schematic representation of osmotically driven and pressure driven membrane processes. Respectively from left to right: forward osmosis (FO), assisted forward osmosis (AFO), pressure retarded osmosis (PRO), and reverse osmosis (RO). The low saline solution is displayed in light blue (this is pure water in case of RO), while the high saline solution is shown in dark blue. This figure was adapted from Cath, et al. [9].
4
1.1. Forward osmosis
1.1.1. Principles Similarly to RO, forward osmosis makes use of a selectively permeable membrane that only allows
water to pass through and rejects dissolved solutes. As already highlighted in the introduction of this
chapter, in RO a hydraulic pressure serves as the driving force for mass transport (of water) through
the membrane. However in FO, it is the osmotic pressure difference across the membrane that is
responsible for water permeation, i.e. water occurs through osmosis. Consequently, water transport
takes place spontaneously from the region of low osmotic pressure, i.e. the feed solution (FS), to the
region of high osmotic pressure, i.e. the draw solution (DS). This is illustrated in Figure 1. Because the
osmotic pressure is solely dependent on the concentration and composition of the solutions,
increasing the osmotic pressure difference will result in a higher driving force and subsequently a
higher water flux [9].
A rather new concept among the ODMPs is assisted forward osmosis (AFO), which is closely related
to FO. Hence, only a few authors have been exploring the viability of the technique [10, 11]. In
contrast to RO and pressure retarded osmosis (PRO) (Paragraph 1.2.), where the hydraulic pressure is
directed in the opposite direction compared to the osmotic pressure, AFO uses the external pressure
as an extra driving force during the water permeation process, in addition to the osmotic pressure
(so in the FO arrangement, this means putting extra pressure on the feed to push more water to the
draw solution). Through this alignment, the water flux will be enhanced resulting in a higher
efficiency of the system, potentially increasing its recovery or lowering the required membrane
surface area [10].
1.1.2. Applications The first practical application of the FO process was carried out by Kessler and Moody [12]. Using a
nutrient solution with a high osmotic pressure, fresh water was extracted from seawater.
Nevertheless, not much research has been conducted in the coming decades, as proper membranes
for FO purposes were desperately required in order for the process to break through. This change
took place as Hydration Technology Inc. (HTI) managed to produce reasonably permeable cellulose
tri-acetate (CTA) membranes during the last decade [13]. Since then, many opportunities appeared
[9, 14-17].
In case of seawater desalination, FO can be operated in different modes. One possibility is creating a
hybrid with RO. Here, the FO process serves as a pre-treatment step, diluting seawater by using
impaired water as the feed. Diluted seawater requires less hydraulic pressure when treated in the RO
installation, resulting in a more energy efficient desalination process [14]. In addition, FO can be used
5
for post-treating the brine of the RO part as well by diluting it with the same water source applied in
the pre-treatment step (Figure 2b) [15]. However, FO can also exist as a stand-alone process,
whereby an artificial draw solution is used (because a highly saline solution is required to draw water
from seawater) which is then separated in a draw solution recovery step after the FO part to produce
potable water. Up to now, the most promising artificial draw is produced by mixing ammonium
carbonate and ammonium hydroxide in specific proportions, giving rise to ammonium bicarbonate,
ammonium carbonate, and ammonium carbamate as salt species. To regenerate this draw solution,
the diluted draw solution can be heated up to 60°C, making it decompose into ammonia and carbon
dioxide and leaving behind purified water. The distillate can then be reconcentrated in order to reuse
it as a draw solution. A major drawback of this system is the energy requirement situated in the
regeneration part, which doesn’t make the system more beneficial than RO unless low-grade waste
heat derived from industrial processes or renewable solar energy is applied [16]. In general for FO
desalination, the lack of high-performance membranes and easily separable draw solutions are
currently limiting the viability of the stand-alone process [9]. In this study, the focus lies on the stand-
alone process as it is simulated on a lab-scale in order to investigate the effect of the high ionic
strength of seawater on membrane fouling. Next to investigating fouling behaviour in this “worst-
case scenario” (regarding the ionic strength), it will enable the evaluation (to a limited extent) of
using novel generations of membranes in this process.
In addition to desalination purposes, FO has many other (potential) applications including: treatment
of industrial wastewaters [17], concentration of landfill leachate, treatment of liquid foods [18],
implementation in life support systems [19], and more.
Figure 2. Schematic representation of the application of FO in seawater desalination by (A) a stand-alone process using an artificial draw solution in a closed-loop system or (B) a hybrid with RO, in which the seawater (brine) is diluted by the FO process. The symbol “O” represents a low pressure pump, which is required for the circulation. These figures were adapted from McCutcheon, et al. [16] and Bamaga, et al. [14] respectively.
6
1.1.3. Advantages As described above, FO bears inherent potential advantages (depending on the application) which
make it attractive compared to the conventional RO process. In terms of seawater desalination,
these advantages include: (1) the absence of a high hydraulic pressure, leading to a lower energy
demand in case no draw solution needs to be separated or the separation is performed using low-
grade waste heat, (2) less sensitive for fouling since compaction of the fouling layer does not occur -
not applying an external hydraulic pressure results in a easily reversible fouling layer (Paragraph 2.6.),
and (3) comparable water quality to the RO process in case of the standalone process, yet after using
a draw solution separation step [9, 20].
1.2. Pressure retarded osmosis
1.2.1. Principles Pressure retarded osmosis (PRO) can be interpreted as an intermediate process between RO and FO.
Similar to RO, a hydraulic pressure is applied, opposite to the osmotic pressure gradient [9].
However, the direction of the water transport is in line with FO, because the applied hydraulic
pressure does not exceed the osmotic pressure difference across the membrane [9]. Illustrated in
Figure 1, water from the feed solution (e.g. fresh water) flows through the membrane into the draw
solution (e.g. seawater), hereby diluting the latter. The resulting diluted and still pressurised draw
solution is then separated into two streams as shown (Figure 3). One of them is relaxed over a
hydropower turbine in order to extract energy to a higher extent than the energy required for
pressurising the draw. The other stream flows through a pressure exchanger to pressurise the
incoming draw solution, further increasing the energy output of the process [21]. As such, energy can
be generated from salinity gradients.
1.2.2. Applications Pioneered by Loeb and co-workers in the seventies [22], PRO has received great renewed interest
during the last decade as well [21, 23-26]. Recent research included the design of specific modules
for the PRO process and the development of CTA FO membranes, which have led to increased power
outputs and eventually the first full-scale prototype PRO installation by Statkraft [27]. Furthermore,
Statkraft has analysed the potential osmotic power exploitation in Norway, Europe, and the entire
world. Respectively, 12 TWh/year, 170 TWh/year, and 1655 TWh/year were the estimated potentials
[26]. However, the pilot installation demonstrated that, in order to be economically efficient, a
power density higher than 5 W/m² is required. This couldn’t be achieved by the early experiments by
Loeb and co-workers (0.1 W/m²) [22] and is still not achieved using the CTA based FO membranes
7
from HTI (3.5 W/m²) [26], currently limiting the feasibility of the PRO process. Therefore, Statkraft
has currently put further experiments to a halt, waiting for future membrane developments [27].
The abovementioned potential power outputs, estimated by Statkraft, were based on the presence
of estuary regions [21]. In those regions, both fresh and salt water are present, making it an ideal
location for a PRO installations. However, it is clear that these areas are limited. Therefore, Gormly,
et al. [28] proposed to increase the potential of coastal regions by locating the PRO set-up next to
waste water treatment plants, which then use the discharged wastewater stream as the feed. This
system could provide the production of sustainable energy - up to 26 MW for California alone [28]- in
combination with a tertiary (advanced) treatment of the wastewater.
Figure 3. Schematic diagram of osmotic power generation. As the water permeates from the feed to the draw solution side, the latter is diluted. One part of this still pressurised draw solution then passes a hydropower turbine while the other part is sent through a pressure exchanger to pressurise the incoming draw stream (seawater). The figure was adapted from Achilli, et al. [23].
In addition to stand-alone configurations, PRO can be used in a hybrid configuration with RO as well,
reducing the energy demand of the RO plant dramatically. Basically, the position of the PRO
installation can be situated either in front of the RO plant, similar to FO (Paragraph 1.1.2.), and/or at
the end of the desalination process. In case of the former, next to energy production also seawater
will be diluted, lowering the energy requirements for the subsequent RO process. For the latter, the
highly concentrated brine solution originating from the RO process is used as the draw stream in the
PRO part, resulting in a high osmotic pressure difference and consequently a high amount of
extracted energy [29, 30]. It should be noted that the RO-PRO approach has only recently developed
from a pure theoretical approach to the first pilot-scale installations [31, 32], indicating that the
applicability of PRO is still in its infancy.
1.3. Mass transfer in osmotically driven membrane processes In the aforementioned sections, the principles regarding pressure driven as well as osmotically driven
processes were explained. To recall: osmosis is driven by a difference in solute concentration across
the membrane, i.e. an osmotic pressure differential ∆ , which allows water to permeate from the
low to the high salinity solution. This requires membranes that allow water to pass, but reject most
8
of the solutes. Transport of water in FO, PRO, AFO, and RO can be described by one general
equation, which is based on the solution-diffusion theory [33]:
= (Eq. 1)
where is the water flux (L/m²/h), A is the water permeability coefficient of the membrane
(L/m²/h/bar), is the osmotic pressure differential across the membrane (bar), and is the
hydraulic pressure differential (bar).
Illustrated by Figure 4, for FO, ∆P = 0; for PRO ∆P < ∆ ; for AFO, ∆P < 0 and for RO, ∆P > ∆ (a
negative value for the pressure is obtained in case its direction is opposite to the water flux and v.v.).
As stated above, PRO has to potential to generate power out of a salinity gradient. In order to
express the generated power per unit membrane area, the power density (W) is defined [33]:
W = (Eq. 2)
Combining Eqs. (1) and (2) gives rise to:
W = (Eq. 3)
The latter equation indicates that
and thus a maximum power density is reached when
the hydraulic pressure applied on the draw is half of the osmotic pressure, which is illustrated by
Figure 4.
Figure 4. Water flux (Jw) and power density (W) versus applied hydraulic pressure (ΔP) comparing ODMPs with RO. Both Jw and ΔP are orientated from the draw to the feed solution. This graph was adapted from Achilli, et al. [23].
9
CHAPTER 2 - FACTORS INFLUENCING THE PERFORMANCE OF
OSMOTICALLY DRIVEN MEMBRANE PROCESSES In the previous chapter, ODMPs were described and compared with a traditional pressure driven
process (i.e. RO). Despite being a promising alternative to conventional methods, there are still areas
of ODMPs that require extensive research before full-scale applications can become reality, including
the availability of efficient draw solutions, the fabrication of (specific) high flux membranes and the
persistent problem of membrane fouling. The latter two are the main areas in which this work is
situated, as those fields are inextricably linked. Several other factors affect the performance as well
as the viability of ODMPs and will be therefore outlined and discussed below.
2.1. Concentration polarisation The water flux in ODMPs was given by Eq. 1. In this equation, the osmotic pressure difference ∆ is
defined across the active layer of the membrane. However, ∆ is much lower than the bulk osmotic
pressure difference, resulting in a much lower water flux [34]. The cause of this smaller osmotic
pressure difference, and thus reduced flux,, is mainly attributed to a phenomenon called
concentration polarisation. The existence of the latter was already observed in research regarding
pressure driven processes [35], and was referred to as (concentrative) external concentration
polarisation (ECP).
In RO, the water from the high salinity solution is pushed through the membrane. This induces a
convective flow, which drags solutes from the bulk solution to the active surface layer of the
membrane. As the solutes are rejected by the active layer of the membrane, the concentration rises
which increases the osmotic pressure difference. Consequently, an even higher pressure is required
to overcome this gradient in order to maintain the same flux (Eq. 1).
Unlike its pressure driven variant, concentration polarisation (CP) is more complex in FO: since salt is
present on both sides of the membrane, CP can occur on both sides of the membrane. If a dense,
symmetric membrane is considered, concentrative external concentration polarisation (CECP) would
occur at the feed side of the membrane (Figure 5a, right side) due to the convective flux, increasing
the osmotic pressure, similar to pressure driven membrane processes. Simultaneously, on the other
side of the membrane the draw solution is being diluted due to the permeating water, lowering the
osmotic pressure of the draw solution. This phenomenon is called dilutive external concentration
polarisation (DECP). The overall result is a reduced osmotic pressure difference across the
membrane.
10
In reality however, dense membranes are not often used due to their low water permeability.
Therefore, asymmetric membranes are applied, consisting of a thin, smooth active layer for rejecting
of solutes and a thicker, porous support layer for mechanical strength. This results in another CP
phenomenon referred to as internal concentration polarisation (ICP). Depending on the position of
the active layer, facing the draw solution (AL-DS or PRO mode) or facing the feed solution (AL-FS or
FO mode), two subdivisions can be considered. In the AL-DS mode (Figure 5b), concentrative internal
concentration polarisation (CICP) occurs as the salt in the feed solution enters the open structure of
the support layer due to the convective water flow. Subsequently, the salt will increase in
concentration because it cannot (easily) penetrate the active surface layer, whereas it can penetrate
the support layer easily. The concentration in the support layer increases much more compared to
ECP, mainly since the diffusion of salts back to the bulk solution is more hindered in the porous
support than in the bulk. When reversing the membrane to acquire the AL-FS mode (Figure 5c),
dilutive internal concentration polarisation (DICP) will occur since the permeated water dilutes the
draw solution [34]. It should be noted that ICP is more severe in the AL-FS mode as the highly
concentrated draw solution then faces the porous support layer, creating a hardly removable
concentration gradient [36].
The occurrence of these phenomena needs to be incorporated in the flux equations to assure a
proper prediction of the effective osmotic driving force. For this, the following equations can be used
in case of the AL-FS configuration (Figure 5c) [34]:
(Eq. 4)
(Eq. 5)
where is the bulk osmotic pressure of the draw solution (bar), is the osmotic pressure of
the draw solution at the membrane’s surface, is the bulk osmotic pressure of the feed solution,
and is the osmotic pressure of the feed solution at the membrane’s surface. Furthermore, two
constants appear in this equations, namely k and K, which are the mass transfer coefficient in the
bulk solution (-) and the solute resistivity for diffusion within the porous support layer (-),
respectively. The former is calculated by
, while the latter is calculated by
[34]. D is
the diffusion coefficient, represents the hydraulic diameter (m), S is the structure factor (µm)
(Paragraph 2.3.), and Sh is the Sherwood number (-), which depends on whether the flow is laminar
or turbulent. In case of a flow through a rectangular channel, the Sherwoord number can be found by
[37]:
11
Laminar flow (Re ≤ 2100) 1.85(
(Eq. 6)
Turbulent flow (Re ˃ 2100) 0.04 (Eq. 7)
where Re is the Reynolds number (-), Sc is the Schmidt number (-), and L is the length of the channel
(m). For the AL-DS configuration, the negative signs in Eqs. 4 and 5 are reversed.
These equations enable the calculation of the solution concentrations at the active layer in order to
determine the effective driving force and ultimately the water flux through the membrane. The latter
can be found by the following equation when applying FO in the AL-FS mode:
(Eq. 8)
where B is the solute permeability coefficient (m/s) (Paragraph 2.3. for the latter). Note that for the
complete elaboration of the overall flux equations, one has to take into account several other factors
in the formulas (e.g. reverse solute diffusion, Paragraph 2.2.) [23]. In addition, accurate
determination of the membrane properties (Paragraph 2..3.) is key in order to predict the permeate
flux correctly when using Eq. 8 [38].
Figure 5. Illustration of osmotic driving force profiles for osmosis for (a) a dense symmetric membrane, (b) an asymmetric membrane with the AL-DS configuration (PRO mode), and (c) an asymmetric membrane AL-FS configuration (FO mode) [34].
12
2.2. Reverse solute diffusion In RO, the solute and water flux are orientated in the same direction, which is a major difference
compared to ODMPs. In the latter, solutes are typically transported in the opposite direction
compared to the water, namely from the draw to the feed solution. This phenomenon is called
reverse solute diffusion (RSD). Its presence is directly related to the osmotic pressure difference
across the membrane, necessary to induce the driving force for water permeation, and the fact that
no membrane is able to form an ideal barrier. The equation below describes the solute flux [39]:
(Eq. 9)
where is the solute flux (moles/m²/h), is the solute concentration of the draw (moles/L),
and is the solute concentration of the feed (moles/L). The higher the draw solution
concentration, the higher the RSD (at a constant feed solution concentration). In PRO and AFO, RSD
can become even more severe as a result of enlarged pores of the membrane’s active layer under
the hydraulic pressure, facilitating an easier transport of solutes [25]. As a result of RSD, reduction of
the osmotic pressure in the DS and elevation of the concentration in the FS will lower the osmotic
pressure difference across the membrane, thereby reducing the performance of ODMPs.
Furthermore, replenishment of the lost draw solution can increase the operational cost of the
process (depending on the cost of the draw solution of course). Another possible disadvantage of the
RSD is the requirement of an additional treatment step of the feed solution in case of a detrimental
draw solution (for example in liquid food processing). Last but not least, RSD can have a drastic effect
on membrane fouling, which is a crucial impact that deserves proper attention later on (Paragraph
2.6.3.) [6, 40-42]. Therefore, it is imperative to design membranes with low RSD, and thus a low
solute permeability.
2.3. Membrane properties The inherent properties of the membrane define its overall performance. In general, asymmetric
membranes are used in ODMPs. These membranes consist of a thin, smooth active layer for rejecting
solutes and a thicker, porous support layer for mechanical strength [34]. Membrane properties such
as water permeability (A) and solute permeability (B) are related to the active layer, while another
parameter called the structure factor (S), describes the porous support layer. Already introduced in
Eq. 1, the water permeability quantifies the extent to which the water is able to flow through the
membrane’s structure. Rearranging Eq. 1 gives rise to:
(Eq. 10)
13
In this formula, is the membrane’s surface (m²), is the volume of the permeated water (L)
and is the time elapsed during the permeation (h).
A second parameter, the solute permeability (B), is related to the incomplete solute rejection of the
membrane. It was seen in Eq. 7 that B affects the RSD proportionally [39]. B itself can be expressed
by the following formula [43]:
(Eq. 11)
where R, indicated by a percentage is the rejection of solutes, which can be easily determined by
[44]:
(Eq. 12)
with the concentration of the permeate (moles/L).
In general, pure product water in combination with a high yield (i.e. a high amount of permeate) is
the objective in all membrane-based filtration systems. As such, a high water permeability and a low
solute permeability are desirable for membranes used in ODMP’s, since this maximizes the
performance of the process. However, due to the asymmetric structure of the membranes, another
intrinsic property which is linked to the porous support layer of the membrane exists: the structure
factor (S) [45]. S can be described as [45]:
(Eq. 13)
where is the porosity (-), x is the thickness (m), and is the tortuosity (-) of the porous support
layer. In practice, S is determined by [43]:
(Eq. 14)
where is the diffusivity of the draw solute (m²/s). A higher value of S indicates a ‘thicker’ porous
layer which results in a higher degree of concentration polarisation. Therefore, a low structure factor
is desirable to reduce the loss of performance. Typical values for A, B, and S in case of traditional as
well as novel generations of membranes are presented in Table 1.
In order to measure the intrinsic parameters A, B, and S of an FO membrane, existing approaches
make use of two separate experiments. A and B are typically measured in RO experiments by
applying a trans-membrane hydraulic pressure. Subsequently, S is determined during an FO
14
experiment using an osmotic driving force [11, 46-49]. Nevertheless, a high variability between the
measuring protocols exist (such as draw and feed solutions concentration, type of solutions, cross-
flow velocity, using additional PRO experiments, etc.), and as a result, A, B, and S factors can vary
widely in literature [38]. To overcome this variation, Cath, et al. [38] recently proposed a standard
methodology for measuring the three parameters. However, this protocol still requires multiple
experiments in different setups. In addition, FO membranes are subjected to high pressures during
the RO tests, which can damage the membrane, or change the values of the intrinsic parameters
[10]. Another shortcoming of the methodology is that it’s based on the notion that transport
parameters are universal and transferable. The latter is unlikely because of the fundamentally
different driving forces between pressure and osmotically driven membrane processes. Therefore,
Tiraferri, et al. [50] presented a methodology consisting of a single FO experiment in order to
characterise A, B, and S. No further details will be discussed here on the methodology itself, as this
would exceed the boundaries of this thesis, yet it is important to note that care needs to be taken
when carrying out FO experiments.
The membrane properties mentioned above mainly depend on the type of materials used for the
membrane preparation and the fabrication process applied. Because of the high relevance for this
thesis, as different traditional and novel membranes will be examined, a separate section is
dedicated to membrane materials and development (Chapter 3).
2.4. Membrane orientation Membranes for ODMPs typically consist of an asymmetric structure, as a “thick” porous layer
provides support for the thin, dense rejection layer. This implies that two orientations are possible,
e.g. the AL-DS and AL-FS mode (Figure 5). In Paragraph 2.1., it was outlined that the effects of CP on
the performance of ODMPs depend on the orientation of the membrane. For instance, ICP is more
pronounced in the AL-DS mode compared to the AL-FS mode [51]. In addition, membrane fouling is
highly affected by the orientation of the membrane as well. Zhao, et al. [52] performed experiments
with organic and inorganic foulants, proposing an overall advice for the selection of the membrane
orientation. In general, the AL-FS mode showed a lower flux decline in both fouling experiments, an
observation also made in previous research [53]. In addition, the AL-FS mode always resulted in
higher water flux recoveries, which were determined by normalising the initial water flux after
rinsing the membrane (using de-ionised water and a cross-flow of 25 cm/s for 0.5 h) by the initial
water flux of the fouling run. This can be explained as follows: in the AL-DS mode - where the porous
support layer is facing the foulant containing solution - foulants are deposited in the porous
microstructure of the support layer. Such a structure makes it more difficult for the foulants to be
15
washed out compared to the AL-FS mode, where foulants are deposited on top of the smooth active
surface layer (which is a dense layer, so the solutes are not able to enter it) [52]. In general, it has
been established that the AL-FS mode should be selected for ODMPs in case of a feed solution with
high fouling tendency (e.g. in membrane bioreactors). On the other hand, the AL-DS mode is more
useful when applying feed solution with a lower fouling tendency (e.g. brackish water desalination),
since it results in higher membrane fluxes due to lower ICP [52].
2.5. Operating conditions In addition to the abovementioned membrane properties (A, B, R, and S), which determine the
occurrence of the concentration polarisation and reverse solute diffusion phenomena, the conditions
in which ODMPs are operated have a great influence on the overall performance as well. These
include i.a. temperature and the type of draw and feed solutions.
2.5.1. Temperature Temperature is related to mass transfer, mineral solubility, membrane fouling, and concentration
polarisation, indicating that it is an important operational condition. Zhao and Zou [54] investigated
the influence of temperature on the performance of an FO desalination process. At higher
temperatures, the water flux appeared to increase, resulting in higher water recoveries. The main
reason was a decrease in viscosity, resulting in higher water permeability of the membrane, and an
increase of osmotic pressure (and thus driving force) due to increased solute mass transfer (higher
diffusion coefficient). Similar observations were made by She, et al. [25] and Xie, et al. [55], as the
determined A-values increased with increasing temperature. In addition, also the B-value increased
due to increased solute mobility, resulting in a higher RSD. Nonetheless, the reverse solute flux
selectivity – which is the ratio of the water flux and the salt flux – and the membrane structure factor
S seemed unaffected by an elevation in temperature. Furthermore, it was found that also rejection of
trace organic compounds and membrane fouling were depended on temperature. For instance,
scaling – a phenomenon that occurs when the salt concentration at the membrane surface exceeds
the maximum solubility – became more severe in case of higher temperatures due to increased
crystallisation [54]. However, this is only important in case of salts with a temperature-depending
solubility (e.g. CaCO3). Rejection of (negatively charged) trace organic compounds, on the contrary,
could be improved at elevated temperatures [55]. One can conclude that temperature has a complex
effect on the performance of membrane processes in real life. However, in lab-scale experiments,
temperature variability is usually kept to a minimum by temperature control, reducing its effect.
16
2.5.2. Feed and draw solutions In the previous paragraph, it was shown that the working temperature is an important parameter
with regard to process performance. Nonetheless, also other factors related to the feed and draw
solutions can have a significant contribution to that matter. As the draw solution (DS) is the crucial
source for the osmotic driving force, a proper selection has to be made in order to find the most
suitable DS for a given application – because as summarised in Chapter 1, many applications exist.
According to Ge, et al. [56], an ideal DS needs to fulfil the following criteria:
- Firstly, a sufficiently high osmotic pressure has to be generated. The osmotic pressure can be
calculated by the Van’t Hoff equation [54]:
(Eq. 15)
where i is the Van’t Hoff factor (-), M is the molarity of the solute which is equal to the ratio of the
number of solute moles (moles) to the volume of the solution (L), R is the gas constant (L.atm/K
moles), and T is the absolute temperature (K). Typically, compounds with a high water solubility (low
molecular weight) and a high degree of dissociation fulfil the first criterion.
- Secondly, the reverse solute flux must be minimal, as this was seen to negatively affect the
performance of ODMPs by reducing the effective osmotic pressure difference (Paragraph 2.2Reverse
solute diffusion). In addition, the feed solution gets contaminated, while costs to replenish the DS
rise because of dilution.
- The third criterion requires easy regeneration of the diluted DS, since regeneration consumes
energy - due to hydraulic (RO, NF, etc.) or thermal processes (membrane distillation, etc.) - and
thereby it increases the overall operation costs. Thermodynamically, this is the reason why closed-
loop FO applications can never be energetically more favourable than RO processes, unless energy
sources such as waste heat, for example, can be used to regenerate the draw solution.
- Lastly, a small molecular weight and low viscosity draw solution is necessary in order to reduce CP.
Both requirements are inversely proportional to the diffusion coefficient of the solute [56].
With respect to these terms, various DSs have been tested, including volatile compounds [16],
nutrient compounds [57], inorganic salts [16], and organic salts [58]. Also synthetic materials, such as
nanoparticles [59] and poly-electrolytes [60], have been used as DSs. Further examination of DS goes
beyond the boundaries of this report, but remains a point of attention in ODMPs.
Next to the draw solution, also the feed solution is able to affect the performance of ODMPs. Here,
desirable feed streams have a low salinity in order to maximise the osmotic pressure difference, as
17
well as a low fouling tendency to avoid loss of performance. In practice, typical feed solutions used in
ODMPs are river water, impaired or pre-treated waste water, industrial waste water, seawater (in
the stand-alone FO process), etc. [61]. In all of these streams, foulants are ubiquitous and can give
rise to an undesirable decrease of performance. A more detailed review related to fouling will be
presented in Paragraph 2.6.
2.6. Fouling Membrane fouling is an inevitable phenomenon that occurs in all membrane-based filtration systems
due to the presence of contaminants in the applied water sources (e.g. seawater, river water, waste
water, etc.). During the fouling process, reversible or irreversible deposition and accumulation of
foulants takes place on the membrane surface. Among these foulants, different categories can be
distinguished: colloidal fouling, biofouling, organic fouling, and inorganic fouling (scaling) [62]. In
general, membrane fouling results in a lower water flux, a shorter membrane life, and the necessity
of using cleaning agents, thus increasing the operational and capital costs of the process [62].
Although fouling is a widespread phenomenon, a large distinction between pressure driven and
osmotically driven membrane processes can be made due to the difference in the main driving force.
As mentioned previously, pressure driven processes like RO use large hydraulic pressures (up to 80
bar) in order to achieve sufficient water permeation. Using such pressures gives rise to a more
compact, denser, and thinner fouling layer compared to ODMPs. Furthermore, the relatively low
water fluxes in ODMPs result in lower fouling rates as well. When increasing the cross-flow during
the fouling run, almost complete recovery of the flux can be achieved in FO, while in RO, increasing
the cross-flow often has limited to no significant effects [6]. This indicates that fouling in ODMPs
tends to be more reversible compared to pressure driven processes. Nevertheless, membrane
fouling is inevitable, even in FO. In addition, fouling behaviour depends on the choice of membrane
orientation (AL-DS mode versus AL-FS mode) [52, 53, 63, 64] and use of a (low) hydraulic pressure
(PRO/AFO versus FO) [61]. In the section below, a summary of the known principles related to fouling
mechanisms in ODMPs will be given, as well as an indication of the existing knowledge gaps.
2.6.1. Effect of physical parameters Referring to Paragraph 2.4., there are two membrane orientations possible in ODMPs: the AL-DS and
AL-FS mode. Since both configurations implicate a different material and/or structure to which the
foulant will contact, a distinction in fouling behaviour can be expected. Mi and Elimelech [53] showed
that bovine serum albumin (BSA) and humic acid (HA) fouling was higher in the AL-DS mode as a
result of a lowered shear force (Figure 7). Shear force to limit fouling is usually induced by the cross-
flow, however its effect on the deposited foulants vanishes in the porous support layer since no
cross-flow is possible there. This explains the difference in flux decline compared to the AL-FS mode,
18
where the shear force is not affected by the membrane structure. Nevertheless, other factors play an
important role in membrane fouling as well. The alginate cake layer for example, was not removed
completely by the high applied cross-flow velocity, resulting in equivalent flux decline in both
configurations, which is illustrated in Figure 7. This research group concluded that alginate fouling
was governed by chemical interactions (i.e. calcium binding) (Paragraph 2.6.2.) instead of
hydrodynamic interactions. Such a statement is in contrast with Motsa, et al. [64], who determined
that, in addition to calcium binding, hydrodynamic forces and membrane surface properties (e.g.
porous structure of the support layer) were all concessive in alginate fouling, indicating that the AL-
FS mode is more resilient to fouling compared to the AL-DS mode. Other research groups, Tang, et al.
[63] and [52], also reported similar differences in fouling propensity between both modes.
Figure 6 Schematic representation of the coupled influence of hydrodynamic forces (y-axis) and intermolecular adhesion (x-axis) on membrane fouling in the presence of alginate, AHA and BSA respectively. The first row illustrates the effects in AL-FS mode, while the lower two represent the situation in AL-DS mode [53].
Next to shear force, another important hydrodynamic interaction influences membrane fouling,
referred to as the permeation drag. Whereas shear force is induced by the cross-flow, permeation
drag is caused by the permeate flux through the membrane, orthogonal to its surface. By varying the
concentration of the draw solution, different initial fluxes can be obtained, which then in turn affect
the permeation drag forces. It can be expected that higher permeation drag forces will lead to more
fouling. In case of BSA for example, more fouling occurred when the initial flux increased because of
19
the denser cake layer formed under the higher drag forces, increasing the resistance (of the fouling
layer) to the permeate flux (Figure 6) [53]. Similar results were obtained by She, et al. [41] and Tang,
et al. [63]. Nevertheless, a certain threshold value in terms of permeate flux has to be achieved in
order to induce significant membrane fouling. This phenomenon, called “the critical flux”, was
observed previously in RO [65], after which Tang, et al. [63] and Zou, et al. [42] confirmed its
existence in FO.
2.6.2. Effect of solution chemistry Another crucial factor that affects membrane fouling is related to the feed solution chemistry. Mi
and Elimelech [53] investigated the role of calcium on the fouling behaviour of several organic
compounds (alginate, BSA and HA). It was reported that in the presence of calcium ions, alginate
fouling was enhanced due to intermolecular bridging between the carboxylic groups and even with
the functional groups on the membrane surface. Consequently, a tighter alginate gel layer on the
surface of the membrane was created. By the use of atomic force measurement (AFM), it became
clear that higher adhesion forces - as a result of this bridging - improved foulant-foulant interactions,
implicating more severe fouling. Furthermore, it was also concluded that the extent of interaction
with calcium varies between the used foulants, which is illustrated by Figure 6. Alginate, for instance,
has a more profound interaction with calcium compared to HA and especially BSA, due to the higher
density of carboxylic acid groups. Zou, et al. [42] also observed a severe flux decline when adding
divalent ions (magnesium) to a feed containing organic matter (algae). Again intermolecular bridging
occurred, although in this case between the algae cells and their extracellular polysaccharides (EPS).
All of these studies thus suggest that divalent ions can result in a drastic flux decline due to the
formation of complexes with certain functional groups.
So far, solution chemistry and foulant-foulant related interactions were discussed in case of the
presence of only a single foulant. In reality however, multiple types of foulants occur in natural
waters. Nevertheless, studies with regard to combined fouling in ODMPs are rare and much
knowledge is still lacking. In this context, Liu and Mi [66] studied the effect of combined organic and
inorganic fouling using alginate and gypsum respectively as model foulants. A synergetic effect was
observed as the flux decline in the combined experiment was higher than the sum of flux declines
during the experiments using the individual foulants. Using a microscope-equipped FO cell enabling
direct observation of the fouling behaviour, alginate was found to increase the size of the gypsum
crystals by acting as a nucleus. In addition, crystallisation kinetics were also accelerated by the
organic foulant. Although less intensely investigated, Boo, et al. [67] also observed a higher flux
decline in case of using both organic as inorganic foulants in the feed.
20
2.6.3. Cake enhanced osmotic pressure and the effect of reverse solute
diffusion Once a fouling layer has been formed, solutes are trapped between the membrane and the cake
layer causing the osmotic pressure at the membrane surface to increase, thereby decreasing the
permeate flux. This phenomenon, called “cake enhanced osmotic pressure (CEOP)”, was first
observed in pressure driven processes by Hoek, et al. [68] and later confirmed by Lee, et al. [69] in
FO. Particularly in the case of colloidal fouling, CEOP is considered to be the main factor that causes a
decrease in permeate flux [6]. Even though the permeate flux decline due to organic fouling is mainly
considered to be governed by the resistance of the fouling layer (Paragraph 2.6.1.), CEOP will also
occur (although the relative importance of both factors is different).
Next to the difference in driving force between pressure driven and osmotically driven processes,
which was already mentioned to influence the fouling phenomenon in the introduction of this
chapter, another main distinction between FO and RO results in a difference in fouling behaviour: the
existence of reverse solute diffusion. The latter occurs uniquely in ODMPs, because in pressure
driven processes only one flux direction is possible due to the high applied hydraulic pressure. The
main consequences of this phenomenon are that also the draw solution is able to affect the fouling
mechanisms in ODMPs directly or induce even more flux decline when a fouling layer is existent.
For example when a fouling layer is existent, RSD can increase the osmotic pressure at the feed side
of the membrane in addition to already trapped solutes originating the feed side itself (which was
called CEOP). Consequently, Lee, et al. [6] adapted the name of this phenomenon to “accelerated
CEOP”. In their research, it was stated that accelerated CEOP was the key mechanism for lowering
the water flux after fouling in FO (instead of the actual fouling layer itself). Experiments were
performed using alginate, HA, BSA and silica colloids. Although similar initial fluxes were obtained
during the RO and FO experiments respectively, the flux decline in FO was shown to be significantly
higher. According to the investigators, this difference could only be attributed to accelerated CEOP,
which occurs only in FO.
Not only can RSD enhance flux decline in case a fouling layer is existent, it can also affect the extent
of membrane fouling itself. Zou, et al. [42] were the first to investigate the direct effect of RSD on
membrane fouling in FO, using algae as a foulant in the feed solution. Despite the higher observed
water flux with MgCl2 compared to NaCl (at the same draw solution concentrations), which has been
seen to be a favourable aspect in FO (Paragraph 2.3.), more fouling occurred when using the former
as intermolecular bridging between Mg2+ and the algae took place (which is not the case for Na+ as
mentioned in Paragraph 2.6.2.). As MgCl2 enhanced membrane fouling due to RSD, the investigators
21
elevated the DS concentration, hypothesizing it would increase RSD and thus fouling. The latter can
be illustrated by Eq. 9: increasing the draw concentration (at equal feed concentrations) will enhance
RSD as a higher value for Js is obtained. Their hypothesis was proven to be true as the results showed
more profound fouling behaviour due to RSD when the DS concentrations were increased.
2.6.4. Effect of membrane materials and modifications In the previous section, it was illustrated that foulant-foulant interactions between similar or
different types of foulants greatly influence the fouling rate. Yet another aspect is of great
importance, namely the foulant-membrane interactions. One research group has performed three
different studies to characterise the foulant-membrane interactions using organic [70], inorganic [71]
and colloidal foulants [72]. In all cases, the difference between a CTA and a modified TFC-RO
membrane (in order to make it more suitable for FO applications) was investigated. Each study
revealed the higher fouling potential of the TFC-RO membrane compared to the CTA membrane
using AFM for characterisation. The presence of highly adhesive sites [70], the negative charge of the
functional groups [71] and the relatively higher roughness of the active layer [72] regarding the TFC
membrane were found to be responsible for the observations. In all cases, the fouling experiments
were performed in Al-FS mode, thus foulant-membrane interactions were only examined for
differences in the active layers. Motsa, et al. [64] examined the difference in foulant-membrane
interactions between the active layer and the support layer of a commercial CTA-W membrane in
case of fouling with alginate at varying electrolyte solutions. Here, it was concluded that for each salt
solution, the support layer was less tolerant to alginate fouling compared to the active layer. Such
conclusions were made by determining the surface free energies of adhesion (between alginate and
the membrane) using the extended Derjaguin-Landau-Verwery-Overbeek (XDLVO) approach, which
indicates the extent of interaction between foulants and a certain surface (e.g. a membrane)
(Paragraph 3.4.)
Even after the fabrication process of membranes, it is possible to alter its fouling resistance by
modifying the membrane surfaces. Although surface modification of RO membranes has been
around for several years [73], Nguyen, et al. [74] were first to introduce such modification in FO by
using poly-amino acid 3-(3,4-Dihydroxyphenyl)-L-alanine (L-DOPA), which is a zwitterionic polymer.
After coating commercial CTA membranes from HTI with L-DOPA for 12 h, 30% less fouling was
observed compared to the uncoated samples, proving the potential of surface modification in
general.
22
2.6.5. Fouling control Membrane fouling is structurally different in FO compared to RO, as no hydraulic pressure is applied
and a lower permeate flux is obtained. Consequently, fouling in FO is often highly reversible, which
was observed by i.a. Mi and Elimelech [70], Lee, et al. [6], and Motsa, et al. [64]. Using alginate to
induce organic fouling, it was reported that a low cross-flow rate of only 8.5 cm/s already recovered
68% of the flux after one hour, while a cross-flow rate of 21 cm/s easily achieved 98% flux recovery
after 15 minutes. This observation indicates that the shear force is the main factor in the physical
removal of the organic fouling layer. In comparison, after fouling the membrane in a RO set-up (at 28
bar) only 70% of the flux could be recovered at the highest cross-flow rate. The high reversibility of
the organic membrane fouling in FO was attributed to the loose and sparse structure of the fouling
layer due to the absence of hydraulic pressure and thus no compaction of the layers occurs. Lee, et
al. [6] confirmed the latter by utilizing alginate, BSA, and HA as organic foulants.
The same research group continued to investigate the possibilities to mitigate membrane fouling
[67]. Again, organic foulants (alginate, BSA, and natural organic matter (NOM)) were chosen to
initiate fouling layer formation. Three different hydraulic cleaning methods were tested. Firstly, high
cross-flow velocities were once again studied, confirming the reduced flux decline at high cross-flow
velocities. Secondly, the effect of turbulence created by feed channel spacers was positive as well.
The mass transfer near the membrane surface improved, which reduced the accumulation of
foulants. Lastly, a pulsed flow was applied, lowering the fouling rate because of three aspects: (1)
increasing the turbulence, (2) increasing the shear force (both (1) and (2) lowered the accumulation
of foulants), and (3) inducing membrane movement, which dislodges foulants from the surface.
The reversibility of membrane fouling, as well as the mitigation possibilities, have made FO a very
interesting option among the membrane-based processes in terms of membrane fouling. Although
fouling is still inevitable, significant advantages can be obtained by not requiring any chemical agents
(which can damage the membrane, increase the operational costs, etc.) [62].
2.6.6. Shortcomings in fouling related research Referring to the aforementioned literature, it is obvious that the majority of publications regarding
fouling behaviour in FO have been focussing on understanding fouling mechanisms in case of single
foulants. Such research has only been the first step in this domain, as in real-life applications, a
complex mixture of different species of foulants will occur. In addition, no research (related to FO)
has ever investigated fouling at the elevated ionic strengths of seawater. Therefore, this research is
aimed to be part of the next step in grasping the underlying fouling mechanisms in FO under
combined fouling conditions and at high ionic strengths.
23
CHAPTER 3 - MEMBRANE DEVELOPMENT IN ODMPs In Chapter 1, it was illustrated that ODMPs have regained great interest during the last decade.
Immense advances in fundamental understanding of material development are the main
contributors to this evolution. General requirements for membranes designed for ODMPs are: (1) a
high rejection dissolved solutes, (2) production of high water fluxes, (3) compatibility with selected
draw solutions, and (4) resistance to mechanical stresses caused by the operating conditions – such
as cross-flow velocities and hydraulic pressure [48, 71-74]. However, the development of appropriate
(commercial) membranes for ODMPs, combining all of the necessary characteristics mentioned
above, is still lagging behind compared to reverse osmosis membranes. Nevertheless, due to the
paramount importance of membranes in the overall process efficiency of ODMPs, a general review
with regard to the current state of membrane design and development is given in the next section.
3.1. Types of membranes The renewed interest in ODMPs was sparked by the development of CTA membranes from HTI.
During the last decade, HTI has long been the only commercial provider of membranes for ODMPs
since the first lab-scale membranes produced by Loeb and Sourirajan in the 1960s [75]. Only recently,
two other players are getting close to hitting the market as well, namely Oasys Inc. [76] and Porifera
Inc. [77]. The hydrophilic nature of the CTA membranes allows properly wetted layers, thereby
reducing ICP and increasing the permeate water flux [78]. In addition, CTA membranes have a lower
fouling potential compared to their more hydrophobic counterparts, e.g. polyamide (PA), polysulfone
(PSf), and polyethersulfone (PES) [70-72, 79]. However, the CTA membranes exhibit rather average
water permeability and salt rejection, currently limiting their possibilities. On top of that, CTA
membranes are vulnerable to changes in pH (only applicable in the range of 4-6 pH), resulting in a
drastic degradation of the polymers over time when the pH varies [48].
On the other hand, conventional thin film composite (TFC) RO membranes, consisting of a PA active
layer and a PSf or PES support layer, achieve high salt rejection, have great mechanical robustness,
and provide chemical stability. Even though the direct implementation in ODPMs is excluded due to
the thick and ICP inducing support layer, resulting in very poor water fluxes [78], TFC membranes are
considered as the most promising type of membrane for ODMPs because of their very thin active
layer, which greatly reduces the structure factor and accordingly CP (Paragraph 2.1. and 2.3.). Still,
because of the essential (operational) differences between RO and FO – RO requires a thick and low-
porous support layer to achieve sufficient mechanical strength in order to withstand the high
hydraulic pressures used (Paragraph 1.1.) – redesigning the support layer (I.e. lowering the thickness
and increasing the porosity) was necessary for ODMP applications since limiting ICP is imperative for
24
the performance of the membranes. Research considering the latter has led to several reports about
successfully produced TFC membranes for ODMPs on lab-scale, particularly during the last 5 years
[47, 48, 80-82] (Table 1).
The abovementioned types of membranes can be implemented in different configurations, of which
flat-sheet membranes are the most readily available. These can be easily stacked in either plate-and-
frame or a spiral-wound modules [9]. The downsides of these kinds of modules are the low packing
density and the lower resistance to pressure [83]. Another option is the use of tubular membranes
(tubes or hollow fibers) which are self-supporting, allow a higher packing density compared to the
flat-sheets and are easier to scale up. However, higher fouling propensity and increased CP
negatively affect these module designs [9, 48, 83-87]
Figure 7. SEM images of the cross-section of TFC (PRO) hollow fiber [83] and flat sheet (FO) membranes [87], respectively left and right.
3.2. Fabrication methods Since the introduction of CA membranes in the 1960s, the basics of the preparation process have
remained the same. In general, fabrication of these membranes is carried out by a process called
phase inversion, or more specifically immersion precipitation. Basically, a polymer is transformed
from a liquid into a solid state during this technique. Firstly, the polymer is dissolved in a suitable
solvent (e.g. acetone) or solvent mixture, which may include additives. Next, this so-called “casting
solution” is spread out upon a supporting layer (e.g. a woven or non-woven fabric) using a casting
knife. Subsequently, the solvent is partially evaporated, followed by the immersion of the cast film in
a coagulation bath containing a nonsolvent (e.g. water). Here, an exchange of polymer occurs
between the solvent and nonsolvent resulting in the precipitation of the polymer. Without going into
further detail, the main parameters affecting the structure of the membranes are: the polymer
concentration, the type of solvent/solvent mixture/nonsolvent, and the use of additives. In addition,
post-treatment of the membranes is possible to obtain desirable characteristics [88].
25
As previously mentioned, TFC membranes designed for ODMPs consist of a thin active layer,
supported by a thicker porous layer, each of a different (polymeric) material. Again, phase inversion
is used, in this case to obtain the support layer (e.g. PSf). On top of this layer, a thin active layer (e.g.
PA) is applied using interfacial polymerisation. This self-sealing and self-terminating technique
involves a polymerisation reaction between two very reactive monomers at the interface of two
immiscible solvents. Firstly, the support layer is soaked in an aqueous solution, containing amines.
After removing the excess solution from the surface (in order to prevent uneven polymerisation), the
membrane is soaked into an organic acid chloride solution. Both monomers react, resulting in a thin,
dense polymeric top layer. Because the support layer is already formed, it is possible to achieve an
ultra thin active layer on top of it, which is the main advantage of TFC membranes compared to CTA.
Another important asset of TFC membranes is the possibility to optimize each layer independently to
obtain the most optimal membrane performance (selectivity, permeability, chemical stability, etc.)
[88].
Besides the optimisation of existing fabrication methods, also (re)new(ed) and promising techniques
have been appeared recently. These include layer-by-layer assembly [89, 90] and the use of
nanofibers for ICP reducing support layers by applying a technique called electrospinning [91, 92].
3.3. Membrane design and development As it is shown in Table 1, major developments have taken place in membrane design and
development for ODMPs. Increasing the flux without compromising on the salt rejecting has been
the main objective in nearly all studies since the renewed interest in ODMPs, as these basic
membrane properties are essential for the economical viability of the processes (higher yields,
reduced membrane surface requirements, ...). Nevertheless, for ODMPs to be implemented in real-
life applications, it is of paramount importance to address the inevitable phenomenon of membrane
fouling. None of the listed novel membranes (Table 1) has been examined under fouling conditions,
thereby overlooking a major asset in the process of membrane development. In addition, no
information can be found regarding the design of foulant resistant/reducing membranes. Currently,
the sole study that has been looking at reducing membrane fouling of FO membranes is Nguyen, et
al. [74], using a pre-treatment step to modify the membrane surface. For these reasons, novel
membranes will be investigated under fouling conditions in order to assess their potential properly.
At the same time, using other materials than the commercial CTA membranes will provide important
and necessary information in terms of foulant-membranes interactions.
26
27
3.4. Characterisation of fouling (propensity) In all of the studies mentioned in Paragraph 2.6., fouling of different membranes (CTA and modified
TFC-RO) was evaluated in terms of flux decline which enabled a proper comparison of the different
foulants. However, next to these experimental observations, more fundamental methods can be
used to support and explain the acquired results. As it was previously cited in 2.6.4., chemical
interactions between the foulant and the membrane surface greatly influence the fouling tendency.
One possible way to predict the magnitude of these interactions is applying the extended Derjaguin-
Landau-Verwery-Overbeek (XDLVO) theory, which determines the membrane and foulant surface
free energy, and as such the interaction energy between the two. The latter is described as the sum
of the apolar Liftshitz-van der Waals (LW) component and the polar Lewis acid-base (AB) component
[94]. Application of the XDLVO approach requires the experimental determination of the surface
energy parameters of the membrane as well as the foulant(s), which can be done by performing
contact angle measurements [95]. The obtained surface energies can be subdivided into surface free
energy of adhesion (interaction between membrane and foulant) and surface free energy of
cohesion (interaction between the foulants). A complete elaboration regarding the application of this
characterisation protocol is presented in Part III, Paragraph 2.1.
Traditionally, the XDLVO theory has been applied for interactions between membranes and colloids.
Brant and Childress [95] used three different types of RO membranes under fouling conditions with
silica and polystyrene colloids. This research group observed positive and negative interaction
energies (for both cohesion and adhesion) when using silica and polystyrene colloids respectively,
indicating more attractive forces in case of the polystyrene. The latter was confirmed by the
experimental results as fouling with polystyrene lowered the permeate flux more significantly.
Nevertheless, the XDLVO approach is also suitable for characterisation of organic fouling, which was
proved by Kim and Hoek [96]. Here, ultrafiltration (UF) membranes were fouled with sodium alginate
and BSA. The thermodynamically stable BSA induced a rapid initial flux decline, followed by a
constant limiting flux, whereas thermodynamically unstable sodium alginate continued to show a flux
decline after the rapid drop in initial flux. The repulsive and attractive forces, between the BSA and
sodium alginate macromolecules respectively, were found to be the cause of these observations.
Not only is the XDLVO approach able to support the fundamental understanding of membrane
fouling, it can also be applied for assessing the fouling propensity of (newly developed) membranes.
For this, several studies used the concept of hydrophilicity and hydrophobicity, which are related to
high and low surface energies respectively. Membranes which are more hydrophilic are considered
to be less prone to fouling [47, 70, 79, 87, 93]. Even though the hydrophiliticy of a membrane is able
28
to give an indication regarding its behaviour under fouling conditions, it is far from sufficient as
foulants have an equal share in the extent of interaction between both. Brant and Childress [95] have
demonstrated this by using the XDLVO approach on three different membranes and three different
colloids. One of the membranes was significantly more hydrophilic, which – if the abovementioned
assumption is followed – should be less prone to fouling. However, as the surface energies of the
colloids were very negative (more negative than the surface energies of the membranes were
positive) it was shown that there was no substantial difference in the extent of fouling between the
membranes, proving the importance of the foulant surface properties in the eventual fouling
propensity of the membranes. In addition, foulant-foulant interactions were showed to dictate
membrane fouling once a fouling layer has been deposited on the membrane surface. In other
words, energies of adhesion (indicating the foulant-membrane interactions) and energies of cohesion
(indicating foulant-foulant interactions) are imperative to assess the fouling propensity of a
membrane. Several studies have been using this approach in research regarding pressure driven
processes [41, 95, 97-99], but only one in FO [64]. The necessity of the use of such proven method
for the assessment of a membrane’s fouling propensity in combination with the lack of research
related to comprehending (combined) fouling mechanisms (especially in ODMPs, Paragraph 2.6.6.)
are the reasons for expanding further on the XDLVO approach in the current study.
29
Part III: Materials and methods
1. EXPERIMENTAL DETERMINATION OF MEMBRANE
PERFORMANCE IN ODMPs This thesis was dedicated to examining membrane performance – under fouling conditions – of
traditional and novel generations of membranes (all commercially available). Two phases could be
distinguished: the first phase was meant to investigate the mechanistic effects of (combined)
membrane fouling, while the second phase compared the performance of both generations of
membranes. Thus, interrelating their performance and acquiring more insight in fouling behaviour
was the ultimate goal.
1.1. Membranes Three types of membranes were used in this study. In the first phase, where fouling behaviour was
examined, commercially available HTI OsMem CTA-ES membranes (Hydration Technology
Innovations, Albany) were applied. These were flat-sheet, asymmetrical membranes in which the CTA
is embedded in a polyester mesh to provide support. This type of membrane also served as the
reference in the second phase of the study, where the performance of novel membranes was
evaluated under fouling conditions. The first novel membrane was obtained from HTI as well, namely
an asymmetrical thin film composite membrane (TFC) with a polyamide active layer. The porous layer
is claimed to be adapted compared to the CTA membrane in order to reduce the effect of
concentration polarisation. The third membrane was received from another (confidential) company.
Here, the advanced porous layer was composed out of carbon nanotubes, which are claimed to be
selective for water molecules only, improving the membrane’s pure water permeability and reducing
ICP. From now on, all three membranes will be referred to as CTA, TFC, and POR. The performance
parameters of each membrane are listed further on in Part IV, Table 3. Once received, the
membranes were stored in deionised water (DI) at 4°C.
1.2. Foulants For mimicking the fouling conditions, two different kinds of foulants were selected in this study.
Alginic acid soldium salt (Sigma-Aldrich, St. Louis, MO),, a.k.a. alginate, served as a model organic
foulant, which has been used in several other studies related to membrane fouling as well [6, 61, 66,
67, 70]. Stock solutions of 2 g/L were prepared by mixing the alginate powder for 24 h after which it
was stored at 4°C. As a second type of foulant, silica colloids (Nissan Chemical Industries LTD., Japan),
were chosen to represent colloidal fouling. One particle size was selected, namely snowtex ST-ZL,
which was reported by the manufacturer to be in the range 70-100 nm. Prior to use, all foulant
30
solutions were shaken for 15 minutes to ensure proper dispersion and interaction with all feed
components.
1.3. Feed solution chemistry Mentioned in the previous paragraph, sodium alginate and silica colloids were the respective model
foulants for organic and colloidal fouling. Both of them are major components in seawater, where
they will be affected by the presence of different kinds of salts. Alginate, for example, is known to
form highly arranged complexes in the presence of divalent cations, causing severe permeate flux
decline [53]. On the other hand, divalent cations can destabilise silica colloids, resulting in a higher
extent of fouling as well [40]. To enhance membrane fouling, relatively high concentrations of
foulants were maintained (200 mg/L sodium alginate and 1 g/L silica colloids). The highly saline
conditions of seawater (0.5 M) were mimicked by preparing a mixture of mono and divalent salts
using NaCl and CaCl2 respectively. Using Eq. 16, a total ionic strength of 0.5 M was obtained
preparing solutions consisting of 0.476 M NaCl and 0.008 M CaCl2.
(Eq. 16)
Here, represents the total ionic strength, the concentration of the ion in solution (M), and the
charge of that ion.
1.4. Experimental set-up and protocol A schematic overview of the experimental set-up, mimicking a stand-alone FO desalination system, is
given in Figure 10. The membranes were placed in a custom-made FO cross-flow module, fabricated
out of polymethyl methacrylate (PMMA). In this module, the two flow channels - through which the
draw and feed streams are able to flow - have dimensions of 25 x 5 x 0.1 cm (with rounded edges). A
polypropylene spacer mesh was added in both channels to create turbulence and mimic real
membrane filtration processes. Several pieces of MasterFlex norprene tubing, interconnected to
each other to provide both streams, while a single Cole Palmer Masterflex peristaltic L/S pump (with
2 heads) was used for recirculating the draw and feed solution from the respective bottles. In one of
the pathways, a flow meter was installed in order to keep the cross-flow velocity at a constant level
throughout the fouling runs. During all of the experiments, the mass of the feed solution was
measured OHAUS Defender Scale 5000 (OHAUS, Germany) and logged on a laptop. As the draw
solution gets diluted by the permeating water originating from the feed solution, a reconcentration
system was applied to keep the osmotic driving force at a constant value. For this, a Buchner flask
was used in which the conductivity of the draw solution (which is related to the salt concentration of
31
a solution) was constantly measured. This data is passed to a laptop, which is in turn connected to a
three way valve.
Figure 8. Schematic representation of the bench-scale experimental set-up which mimics a stand-alone FO desalination
system..
This loop enables the use of a threshold value that can be used to adjust that valve. If the
conductivity drops below the threshold value, the draw solution stream is sent through a funnel filled
with salt until the conductivity exceeds the threshold again. Above the threshold value, the draw
solution will be sent directly into the Buchner flask, without being reconcentrated. In this way, the
concentration of the draw solution is kept at a constant level. The excess draw solution – which is a
result of the incoming permeate flux due to the osmotic gradient - is removed by the overflow pipe
of the Buchner flask.
Table 1. Overview different sets of experiments fouling experiments performed on each type of membrane. In case of
alginate, experiments were performed in both the FO and PRO mode.
Experiment Set A Set B
Single and combined Sequential
1 alginate alginate - silica colloids
2 silica colloids silica colloids -alginate
3 combined combined- combined
32
Each set of fouling experiments is presented in the Table 2. For the first phase, the AL-DS mode was
selected to accelerate membrane fouling - as it was seen that the porous support layer is more
susceptible to fouling – which is favourable for investigating factors that govern membrane fouling.
In reality however, the porous layer is meant to support the active layer which in turn blocks salts
and foulants, making it unsuitable for heavily fouled streams. Thus, in order to compare fouling
behaviour of different membranes properly, which is the case in the second phase, the AL-FS mode
was selected. For the same reason, the draw solution concentration was adjusted to obtain an equal
initial flux no matter what membrane is applied. Each filtration experiment was performed for 24 h
with an initial flux of 16.84 ± 0.71 L/m²/h, while maintaining a cross-flow velocity of 30 L/h.
1.5. Determination of water flux – Performance evaluation Water permeation occurs due to the difference in osmotic pressure between the draw and the feed
solution. Determination of the water flux was done by measuring the mass of the feed solution,
which was logged every 5 minutes by the scale. For a given time interval of Δt = , the water
flux ( ) can be calculated as follows:
(Eq. 17)
where is the surface of the membrane (m²), is the mass of the feed solution (kg), t is the time
(h) and is the density of the water (which is assumed to be constantly 1000 kg/m³). is equal to
0.01244 m², as both flow channels are identical.
Prior to each fouling experiment, a baseline test was performed. Here, conditions are identical to the
fouling test except for the presence of foulants. Comparing both experiments will provide insight in
the extent of fouling. Furthermore, normalising all fluxes (indicated as Jw/J0) obtained from both the
filtration experiments allows a proper evaluation of the differences in flux decline between the
different conditions (e.g. different foulants, combinations, and sequences).
33
2. CHARACTERISATION OF MEMBRANES AND FOULANTS Besides the examination of membrane performance in terms of fouling behaviour, a more
fundamental approach was applied to explain the experimental observations, generally known as
characterisation of membrane fouling. Several methods can be applied to do so, as it was already
outlined in Part II, Paragraph 2.6. (AFM) and 3.4. (XDLVO approach). In this study, it was chosen to
apply the XDLVO approach..
2.1. Surface tension measurements using contact angles
2.1.1. Theoretical background As mentioned in Part II, Paragraph 3.4., in order to apply the XDLVO approach, the surface energy
parameters of the membrane and the foulants have to be determined experimentally, generally by
performing contact angle measurements. Implementing the contact angle data into the extended
Young-Dupré equation (Eq. 18), which relates the contact angle (θ) of a liquid on a (flat) solid surface
with the surface tension of the liquid and the solid, the surface tension parameters of a solid surface
can be determined. According van Oss [100]:
(Eq. 18)
where θ is the contact angle, is the Liftshitz-van der Waals free energy component, is the
electron-acceptor component, and is the electron-donor component. The subscripts s and l
correspond to the solid surface and the liquid, respectively. Using three well-characterised probe
liquids with known surface tension components (
,and ), namely deionised water,
glycerol, and diiodomethane, it is possible to obtain the surface tension components of a solid
surface (
, and ).
The calculated surface tensions for membranes and foulants can be used to determine the total
interfacial free energy of adhesion ( and cohesion
per unit area. Eq. 19 represents
the total interfacial free energy of adhesion as the sum of the Lifshitz-van der Waals (LW) and acid-
base (AB) components [101]:
(Eq. 19)
(Eq. 20)
34
(Eq. 21)
where 1 and 2 represent the different surfaces, while 3 indicates the liquid medium. In case of equal
surfaces 1 and 2, the total interfacial free energy of cohesion can be obtained (i.e. ).
expresses the attraction or repulsion between foulants and already deposited particles as well as the
particle stability. On the other hand, describes the attraction or repulsion between two
different solid materials (1 and 2). In both cases, the total interfacial free energies are determined in
a liquid medium (3) [101].
2.1.2. Measurement protocol From Eqs. 20 and 21, it is clear that different surface tension components of both the membrane and
foulants are required to calculate the free energies of adhesion and cohesion. To accomplish this,
samples were taken from each membrane. For the foulants, a separate experiment was required to
obtain a uniform and thick fouling layer that can be subjected to contact angle measurements using
a dead-end cell.. Here, a nanofiltration membrane (NF90) was fouled at 8 bar with alginate, silica
colloids, and their combination. Subsequently, pieces of the osmotic membranes and fouled NF
samples are placed on a cover glass, on which a filter paper was added with roughly the same size as
the membrane sample, allowing to perform hydrated contact angle measurements. In this way, the
effect of water evaporation, which disturbs the accuracy of the contact angle measurements, is
limited. In addition, the liquid-membrane interface during the filtration process can be simulated. In
this work, a computerised Krüss DSA 10-Mk2 contact angle measurement device (Germany) was used
to carry out the contact angle measurements (Appendix A.). For each membrane sample, a minimum
of 10 drops of all three probe liquids were measured, using the sessile drop method. A maximum
measuring time of 0.5 h for each sample was maintained to prevent dissimilarities due to
evaporation. Averaged values were used to determine the final contact angles. The abovementioned
protocol is similar to Motsa, et al. [64], who indicated the importance of using the hydrated method.
35
Part IV: Results and discussion
1. CHARACTERISATION OF FORWARD OSMOSIS
MEMBRANES
1.1. Determination of the membrane performance parameters Three different types of FO membranes have been examined throughout this research. As stated in
Part II, Paragraph 2.3., FO membranes are typically characterised through the determination of their
performance parameters such as pure water permeability (A), solute permeability (B), solute
rejection (R), and sometimes the structure factor (S). After experimental determination of the
membrane permeate flux (Jw) and R, the parameters A, B, and S were calculated for each membrane
following the procedure outlined in Appendix B. For each membrane, the properties are depicted in
Table 3.
For the CTA and TFC membranes, most membrane performance parameters as determined in this
study were similar to those reported in literature [11, 39]. However, the value of S was substantially
higher than typical values reported in literature, which could be a consequence of variations in the
measurement protocol (which was mentioned in Part II, Paragraph 2.3.). Nevertheless, the structure
factor of the TFC membrane was reported to be twice that of the CTA membrane [11], similarly to
the trend observed in this work. The POR membrane was very recently developed and received
under a confidentiality agreement; as such, no other reports exist regarding the membrane
characteristics. The results indicate that, compared to CTA and TFC, the POR membrane exhibits a
large increase in pure water permeability, without an increase the solute permeability. In Part II,
Paragraph 2.1., it was outlined that concentration polarisation (of which the magnitude is related to
the extent of S) is one of the most performance-limiting factors in ODMPs, therefore the most
explicit difference between all types of membranes was the value of the structure factor of POR (344
µm), which is almost half that of CTA membrane and only a fourth of the TFC. Based on these
parameters, it could be hypothesized that currently, POR outperforms CTA and TFC in terms of
membrane characteristics. Further on in this study, also the fouling propensity of the different
membrane types was examined and compared, allowing a more complete assessment in terms of
the overall membrane performance (i.e. membrane properties and fouling propensity combined).
1.2. Determination of the membrane hydrophilicity The next important membrane parameter after solute rejection, structural factor, water and solute
permeability is the membrane’s hydrophilicity (Table 3). This parameter reflects the extent in which
36
the membrane has an affinity for water. In general, a membrane is designated to be hydrophilic
when the contact angle is lower than 30° and hydrophobic if the angle exceeds 90° [102]. The TFC
membrane clearly falls under the 30°-rule, whereas the CTA and POR membrane are less hydrophilic.
In case of CTA and TFC, the measured contact angles are in accordance with literature [11, 64].
Table 2. Overview of the membrane performance parameters and hydrophilicity for each membrane type.
Membrane A B R S Contact angle (water)
L/m²/h/bar 10-7 m/s % µm °
CTA 0.61 1.5 88.5 663 43.1*/67.8**
TFC 1.17 0.2 98.2 1227 25.2
POR 1.89 1.3 96.0 344 44.0 * Active layer ** Support layer
37
2. DETERMINATION OF MECHANISMS GOVERNING
COMBINED ORGANIC AND COLLOIDAL FOULING IN
FORWARD OSMOSIS The first part of this study consisted of determining the factors that govern combined membrane
fouling with alginate and silica colloids at seawater level high ionic strength. As previously mentioned
in Part III, Paragraph 1.1., a standard CTA membrane was applied in AL-DS mode. Using this
configuration was meant to enhance membrane fouling, making it easier to elucidate the membrane
fouling mechanisms.
2.1. Single and combined membrane fouling Figures 11 and 12 display the permeate flux decline trends for the different fouling runs and the
baseline experiments (i.e., an experiment at similar ionic strength, but lacking foulants) in the
absence and presence of salts, respectively. It can be observed that the final flux decline of the
baseline in Figure 11 is approximately 10% lower compared to the baseline presented in Figure 11,
which is a direct consequence of the limited feed volumes used in each experiment. During the
filtration experiments, water permeated from the feed solution to the draw solution, while salts
were being rejected. Thus, when salts are present in the feed solution (which is the case in the
experiments depicted in Figure 12), an increase in salt concentration is caused by the water
permeation to the draw solution, resulting in a (limited) flux decline during the baseline experiment.
In the absence of salts in the feed solution, nearly identical flux decline trends appeared (Figure 11).
During each fouling run (alginate (ALG), silica colloids (ST-ZL), and their combination (Comb),
respectively), no substantial flux loss could be observed compared to the baseline experiment,
indicating that no (or at least only in a very low extent) fouling occurred. In case of alginate, this
corroborates previous literature [53], where only 10-15% permeate flux loss was observed after 24 h
in an experiment that was conducted with an initial flux that was around 25% higher compared to
this study (21.35 L/m²/h versus 16.71 L/m²/h) but otherwise identical experimental conditions.
Interestingly, results obtained here contrast results that were reported in literature for combined
fouling with alginate and silica colloids. Kim, et al. [103] indicated that upon combined fouling of
alginate and silica, a steady permeate flux decline could be seen from the start of the experiment up
to 30% after already 5 h. Kim, et al. [103] also observed synergetic effects – indicating that the
combined fouling was worse than the sum of the individual foulants – which was not the case in the
work performed in this study. This is surprising, since similar membranes and foulant concentrations
were used. However, the membrane was applied in AL-FS mode, not in AL-DS. Typically, the AL-FS
38
mode is less susceptible to fouling, but since it is the only difference between our study and theirs, it
could be a possible explanation for the observed dissimilarity.
When the ionic strength of the feed solution was raised to 0.5 M (similar to seawater), completely
different trends were seen, and substantial differences in fouling patterns occurred between the
single and combined flux decline trends (Figure 12). This is mainly due to the presence of calcium
ions, since the calcium was previously reported to increase the fouling potential of alginate [53, 104]
and silica colloids [40] as single foulants, due to intermolecular bridging and colloid destabilisation,
respectively. Only after 8 h of fouling, the flux decline curve of the silica colloids started to deviate
strongly from the baseline, indicating that either the silica colloids required some time to settle on
the membrane surface, or it takes time before a sufficiently thick (and perhaps compact) cake layer is
formed that causes the permeate flux to drop. In contrast, alginate fouling depicted a more rapid
initial fouling rate (0.0948 h-1, Table 4) after which the flux decline gradually became more stable.
Interestingly, the combination of both foulants showed an initial flux decline rate (0.1010 h-1) and
flux decline trend similar to that of alginate fouling alone, even though the total concentration of
foulants is much higher in the combined case (0.2 g/L versus 1.2 g/L). This is contra-intuitive as a
higher concentration of foulants is expected to result in a higher degree of fouling, at least in case of
fouling with individual foulants. Possibly, at the high ionic strength, alginate fouling (due to
complexation with Ca2+) overwhelms the presence of silica colloids.
To further examine the effect of Ca2+ during combined fouling, a separate experiment was performed
with a 0.5 M total ionic strength feed solution in absence of CaCl2 (Figure 13). Only when 0.008 M
CaCl2 was present in the feed, a clear drop in permeate flux could be observed. This observation
together with the rather low flux decline observed during the first 10 h of individual silica fouling
(Figure 12) suggest that the fouling behaviour of the combined case is mainly governed by alginate
and its complexation with Ca2+. Furthermore, these results suggest that no clear evidence can be
found of synergetic effects between both foulants at the given concentrations of mono and divalent
salts, which clearly contradicts the findings made by Kim, et al. [103], who reported clear synergetic
effects with and without the presence of cationic species (although at a lower ionic strength).
A last observation that can be derived from Figure 12 concerns the flux decline trend at the end of
the fouling runs. Here, a substantial deviation from the baseline can be observed for each foulant,
indicating that the effect of the aforementioned loss in osmotic driving force is not dominant. Thus,
the strong deviation from the baseline is most certainly a consequence of membrane fouling.
Moreover, it seems that two distinctly different factors are at play, which is suggested by the course
of the flux decline trends in the final hours of the experiments. In case of alginate and combined
39
fouling, the flux decline trends have reached approximately a constant final flux value, after which no
further membrane fouling occurred. Interestingly, the silica colloids showed no sign of a decrease in
flux decline rate towards a constant flux value after the same time period. These results suggest that
alginate as well as combined fouling are mainly governed by the hydraulic resistance of the fouling
layer to the permeate flux, resulting in a so-called “limiting flux”. This phenomenon implies that no
(further) membrane fouling occurs once the permeate flux drops below a certain threshold value.
From that point, the permeation drag is too weak to ensure sufficient deposition of foulants on the
membrane surface, which is typically indicated by the flux decline trend as it reaches a plateau
(Figure 12) [41, 63]. On the other hand, cake enhanced osmotic pressure – which causes a decrease
in the osmotic pressure difference as solutes get trapped between the membrane and the fouling
layer (Part II, Paragraph 2.6.3.) – is generally considered to be the main factor that governs flux
decline in case of colloidal fouling [6, 68, 69]. Since the flux decline trend in case of fouling with silica
colloids shows a continuous drop (without reaching a plateau), this might indicate that indeed CEOP
is the main cause for the flux decline.
However, more experiments were conducted in the next section in an attempt to prove the
difference in fouling mechanisms (i.e. hydraulic resistance and CEOP).
Time (h)
0 5 10 15 20 25
No
rmal
ised
flu
x
0,0
0,2
0,4
0,6
0,8
1,0
ALG
ST-ZL
Combined
Baseline
Figure 9. Permeate flux decline curves in case of fouling a CTA membrane with single foulants (1 g/L silica colloids (ST-ZL) and 200 mg/L alginate (ALG)) and their combination (Comb) in the absence of mono and divalent salts. The draw solution was adjusted to give a similar initial flux to that of fouling in the presence of background electrolytes. Each fouling experiments was performed in AL-DS mode at a cross-flow of 30 l/h for 24 h.
40
Time (h)
0 5 10 15 20 25
No
rmal
ised
flu
x
0,0
0,2
0,4
0,6
0,8
1,0
ALG
ST-ZL
Combined
Baseline
Figure 10. Permeate flux decline curves in case of fouling a CTA membrane with single foulants (1 g/L silica colloids and 200 mg/L alginate) and their combination in the presence of mono and divalent salts (0.476 M NaCl and 0.008 M CaCl2). Each fouling experiment was performed in AL-DS mode at a cross-flow of 30 l/h for 24 h.
Time (h)
0 2 4 6 8 10
No
rmal
ised
flu
x
0,0
0,2
0,4
0,6
0,8
1,0
0.5 M NaCl
0.476 M NaCl + 0.008 M CaCl2
Baseline
Figure 11. Permeate flux decline curves in case of combined fouling a CTA membrane (200 mg/L alginate and 1 g/L silica colloids) with and without the 0.008 M CaCl2. Each fouling experiment was performed in AL-DS mode at a cross-flow of 30 l/h for 10 h.
41
2.2. Consecutive fouling After the single and combined fouling runs, a new set of experiments was performed. Here, in order
to understand the possible inter-foulant interactions during combined organic and colloidal fouling,
the membrane was fouled with the different individual foulants in alternating sequences. In addition,
the relative influence of hindered back diffusion and hydraulic resistance is further examined as well.
This experimental procedure - performed for the first time in this study - is an entirely new way of
investigating membrane fouling.
As the first fouling run is an exact copy of the experiments presented in the previous section
(Paragraph 2.1.), the discussion of this part will focus on the fouling trends depicted by the second
fouling run after the membrane surface has been modified by the first foulant. The first fouling runs
are thus a duplication of the former experiments, proving the reproducibility of the fouling
experiments.
According to Figure 14, when a fresh feed solution with an ionic strength of 0.5 M was introduced at
the start of the second fouling run, a very distinct difference in flux decline patterns appeared
compared to the flux decline trends of the foulants in the first run. Clearly, two observations can be
made when looking at the consecutive run: (1) a partial restoration of osmotic driving force, and (2)
differences in the structural properties of the fouling layer originating from the first run.
(1) A partial recovery of the initial flux can be either due to an increase of osmotic pressure, occurring
when the new feed is applied (24 h in the first fouling run, the ionic strength of the feed water will
have increased due to water permeating to the draw side and salt being rejected) or a disturbance of
the existing fouling layer caused by introducing the fresh feed. No major differences in the amount of
flux recovery were observed between the different foulants, indicating that the flux recovery is
mainly due to the increase in osmotic driving force.
(2) A distinct difference occured in the first hours after this point of partial flux recovery (which is
marked with the dotted line in Figure 14): the flux decline trends of the second runs regarding the
ALG – ST-ZL and Comb – Comb sequences exhibited a rather high initial flux decline rate (0.0453 h-1
and 0.0994 h-1, respectively) towards approximately the same final flux decline percentage obtained
in the end of the first run. In contrast, the ST-ZL – ALG sequence showed a low and steady flux
decline trend after the flux was partially recovered. These observations support hypothesis proposed
in the end of Paragraph 2.1., stating that seemingly different fouling mechanisms govern fouling with
silica colloids on one hand and alginate and combined fouling on the other hand. It is believed that
the high flux decline rate, thereby reaching the final flux decline percentage in a short time period, is
42
due to the high hydraulic resistance of the existing fouling layers (thus originating from the first run,
which is alginate and combined fouling in this case). In contrast, the slow initial flux decline rate
(0.0122 h-1) during the second run of the ST-ZL – ALG sequence suggest that the CEOP is lowered due
to the diluting effect of a fresh and less concentrated feed.
Table 3. Initial flux decline rates for each run during the sequential fouling experiments.
Sequence Flux decline rate (h-1) Final flux decline (%)
First run Second run First run Second run
ALG - ST-ZL 0.0948 0.0122 65 62
ST-ZL -ALG 0.0173 0.0453 67 68
Comb - Comb 0.1010 0.0994 71 73
Time (h)
0 10 20 30 40
No
rmal
ised
flu
x
0,0
0,2
0,4
0,6
0,8
1,0
ALG - ST-ZL
ST-ZL - ALG
Comb - Comb
New feed
Figure 12. Flux decline curves in case of layer by layer membrane fouling of CTA membranes with single foulants– in alternating sequences – and a combination of both foulants. A fresh feed solution with a total ionic strength of 0.5 M was used for the subsequent fouling run. Each fouling experiment was performed in AL-DS mode at a cross-flow of 30 l/h for 24 h.
43
Time (h)
0 10 20 30 40
No
rmal
ised
flu
x
0,0
0,2
0,4
0,6
0,8
1,0
ALG - ST-ZL
ST-ZL - ALG
Comb - Comb
New feed
New feed
Figure 13. Flux decline curves in case of layer by layer membrane fouling of CTA membranes with single foulants – in alternating sequences – and a combination of both foulants. For the subsequent fouling run, the total ionic strength of the feed solution was adjusted to match the total ionic strength in the end of the previous run. Each fouling experiment was performed in AL-DS mode at a cross-flow of 30 l/h for 24 h.
In figure 14, a partial flux recovery was seen at the point where the new feed solution is introduced.
Most likely, this is due to a lowering of the feed osmotic pressure by introduction of the new feed (as
the feed from the first run at this point is already concentrated). To investigate whether the partial
flux recovery is only caused by this osmotic effect, a new experiment was conducted in which the
osmotic pressure of the new feed that was introduced in the second fouling run, was corrected to
the increased osmotic pressure at the end of the first fouling run. If the partial flux recovery is only
an osmotic effect, this adaptation of the fouling protocol would have to get rid of this partial flux
recovery.
Figure 15 – by eliminating the osmotic pressure effect – clearly clarifies the kind of interactions
occurring between the alginate, the silica colloids, and the membrane as the initial permeate flux was
partially recovered only in case of the ST-ZL – ALG sequence. From this, it can be hypothesized that
indeed the silica colloids are only loosely bound to the membrane (and to themselves) when they are
present as single foulant. When the alginate is introduced in the second fouling run, due to the
adhesion of alginate to the silica colloids, a shearing effect is induced and the colloids are dragged
away from the surface by the alginate, thereby effectively reducing the colloid cake layer thickness
and recovering a part of the permeate flux. Alginate binds more strongly to the membrane, and also
44
when the colloids are present together with alginate the fouling layer is attaching more strongly, as
no partial flux recovery is seen in these cases.
In order to demonstrate the shearing effect of alginate on the silica colloids, a supplementary
experiment was performed in which silica colloids were introduced in the first and the second fouling
run, also adjusting the osmotic pressure of the fresh feed (Figure 16). In this case, no permeate flux
was recovered since the colloids present in the feed solution did not shear away the initially
deposited colloids. Instead, there was a smooth continuation of the first flux decline trend. This
observation further reinforces the hypothesis on the coupled adhesion and shear effect of alginate
on the silica cake layer.
Time (h)
0 10 20 30 40
No
rmal
ised
flu
x
0,0
0,2
0,4
0,6
0,8
1,0
ST-ZL - ST-ZL
New feed
Figure 14. Flux decline curves in case of consecutive membrane fouling of CTA membranes with silica colloids. For the subsequent fouling run, the total ionic strength of the feed solution was adjusted to match the total ionic strength in the end of the previous run. Each fouling experiment was performed in AL-DS mode at a cross-flow of 30 l/h for 24 h.
To recapitulate: the set consecutive fouling experiments has indicated that alginate adsorbs on the
silica colloids (at the high ionic strength). Consequently, it can be expected that due to this
adsorption, silica colloids are embedded within the alginate gel layer during combined fouling.
Nevertheless, it was believed that the alginate complexation with Ca2+ - resulting in a cake layer with
a higher hydraulic resistance - is the main factor governing flux decline during combined fouling
(despite the high concentration of silica colloids in the feed), whereas flux decline was mainly
attributed to CEOP in case of individual fouling with silica colloids. In addition, it seems that the silica
colloids do not affect the overall resistance of the fouling layer, since no major difference in flux
45
decline trends occurs between fouling with alginate and the combination of both foulants (Figure
12).
2.3. Foulant-foulant and foulant-membrane interaction energies
2.3.1. Interfacial free energy of cohesion, ΔG131 Table 5 and Table 7 respectively, show the surface free energies and the subsequent interfacial free
energies of cohesion (ΔG131) for each foulant, calculated using Eqs. 13-15 and the data obtained from
contact angle measurements with the different probe liquids (Table 5). ΔG131 describes the energetic
favorability of a solid material interacting with itself in a liquid medium [95]. In this case ΔG131 will
offer insight on the extent of interactions between foulant particles (of the same material) in our
feed stream. More specifically, negative energies indicate that the foulant is more attracted to itself
than to water, more or less indicating hydrophobicity. Thus, negative interaction energies also
indicate that once a membrane is covered with a certain type of foulant particles, the foulants (of the
same material) in the bulk solution will be attracted to that fouling layer.
The interfacial free energies of adhesion behave in the following trend: ST-ZL (24.19 mJ/m²) > Comb
(-2.51 mJ/m²) > ALG (-22.07 mJ/m²). It is clear that alginate is strongly attracted to itself in water (at
the high ionic strength used), while the silica colloids repel each other. The cohesive energy of their
combination can be found in between that of both foulants. Although it is slightly negative, there is
no clear tendency for attraction or repulsion.
2.3.2. Interfacial free energy of adhesion, ΔG132 The foulant-membrane interfacial free energy of adhesion (ΔG132), in this case describing the
interaction between different foulants and the clean membrane surface, provides insight in the
likelihood of a foulant being attracted or repelled by the clean membrane [96]. Table 7 displays the
calculated foulant-membrane interfacial free energies of adhesion, which were computed using Eqs.
13-15 and the data of Table 6. The adhesion energies showed a similar trend compared to the
cohesive energies for the foulants: ST-ZL (16,74 mJ/m²) > Comb (-2.81 mJ/m²) > ALG (-20.16 mJ/m²).
Only in case of single fouling with alginate and combined fouling, negative values were obtained,
suggesting the possibility for spontaneous adhesion of the foulants on the membrane surface. In
contrast, due to the positive adhesion energies of the silica colloids and the membrane (and thus the
repulsion between them), it could be expected that no immediate silica fouling will take place unless
a sufficiently large driving force is present that can overcome the repulsion.
46
It should be noted that care has to be taken when interpreting the free energies of adhesion as well
as the free energies of cohesion in case the combined foulants, since no clear evidence was found
that when alginate and silica colloids are combined, the foulants behave as one.
2.3.3. Relation between interfacial free energies and membrane fouling As the permeate flux induces a hydrodynamic drag force, any foulant particle (either alginate, silica
colloids or their combination) is pushed to the membranes surface. However, at the surface, binding
of foulant particles onto the membrane is governed by chemical interactions. Consequently, foulant-
membrane interaction energies are expected to control the initial fouling behaviour. Once the
foulant particles cover the entire membrane surface, inter-foulant interactions become more
prominent compared to foulant membrane interactions. At that point, interaction energies between
approaching foulant particles and already deposited foulant particles (i.e. free energies of cohesion)
are expected to dictate further (long-term) membrane fouling [64]. The initial flux decline rates
(calculated for the first 2 h) and later flux decline rates (calculated between 8 and 10 h) were
correlated to the free energy of adhesion and free energy of cohesion in case of CTA (when applied in
AL-DS mode), respectively. A reasonable correlation exists between the initial flux decline rates and
the free energies of adhesion with a correlation factor of 0.711. On the other hand, a rather weak
correlation can be found between the later flux decline rates and the free energies of cohesion with
a correlation factor of 0.617. To some extent, it is shown that no matter what foulant was applied
(either alginate, silica colloids or their combination), initial fouling behaviour and long-term
membrane fouling are mainly governed by foulant-membrane and inter-foulant interactions,
respectively, which is in accordance with the fouling behaviour depicted in Figure 12. Here, silica
colloids were shown to foul the membrane to a lower extent compared to alginate and their
combination (during the first half of the experiment), caused by the repulsive forces between the
colloids and the membrane, and between the colloids themselves. In contrast alginate fouled more
severely, which is a consequence of the observed attractive forces between both the alginate
aggregates and the membrane, and between alginate aggregates. When combining alginate and
silica colloids, the obtained free energies adhesion and cohesion were slightly negative, thereby
reducing the strongly positive energies of the silica colloids to a large extent, indicating the dominant
effect of alginate (which was also stated in Paragraph 2.1.). However, during the fouling experiments
(Figure 12), a flux decline trend occurred that was very similar to that of individual alginate fouling,
suggesting that the XDLVO approach gave rise to a underestimation of the interaction occurring in
reality. This was already indicated in Paragraph 2.3.2.
In addition to the separate fouling experiments, it is also noticeable that the XDLVO predictions agree
to a certain extent with the consecutive fouling experiments. Negative adhesion forces were
47
obtained between alginate and silica colloids (-3.47 mJ/m²), thus supporting the observation that
alginate aggregates tend to adsorb on the colloids, thereby inducing a shear effect on the colloids
cake layer originating from the first fouling run (Figure 15). On the other hand, the strongly positive
cohesive energies (implying repulsive forces) in case of the silica colloids (24 mJ/m²) support why no
shear effect was noticed when performing sequential experiments with only the colloids (Figure 16).
Table 4. Contact angles ( of H2O + 0.476 M NaCl + 0.008M CaCl2) and surface free energies for the single foulants and their combination.
Foulant Contact angle
γLW γ+ γ- γtot (°)
ALG 53.8 38.26 5.16 12.09 54.06
ST-ZL 11.2 36.08 2.68 50.20 59.32
Comb 32.5 48.34 1.62 29.24 62.14
Table 5. angles (H2O + 0.476 M NaCl + 0.008M CaCl2) and surface free energies for the different membranes.
Membrane Contact angle
γLW γ+ γ- γtot (°)
CTA-PRO 67.8 36.56 1E-06 25.30 36.57
CTA -FO 43.1 38.20 0.31 38.50 45.15
TFC 25.2 38.39 3.01 41.49 60.76
POR 44.0 21.84 4.96 34.14 47.88
Table 6. Foulant-foulant interfacial free energies of cohesion and foulant-membrane/foulant-foulant interfacial free energies of adhesion. CTA-PRO refers to the case were the CTA membrane was applied in AL-DS mode, while CTA-FO indicates the membrane was applied in AL-FS mode.
Foulant Energy of cohesion, ΔG131 (mJ/m²) Energy of adhesion, ΔG132 (mJ/m²)
CTA-PRO CTA-FO TFC POR ALG - STZL
ALG -22.07 -20.16 -12.29 -7.32 -4.48 -3.47
ST-ZL 24.19 16.74 22.11 18.90 16.89
Comb -2.51 -2.82 5.03 5.90 7.99 N.A.
48
49
3. COMPARISON OF PERFORMANCE AND FOULING
BEHAVIOUR OF TRADITIONAL AND NOVEL FORWARD
OSMOSIS MEMBRANES A second part of this study was aimed at comparing fouling behaviour of the commercially available
CTA membrane and novel generations of FO membranes (TFC and POR). All fouling tests were
performed in the AL-FS configuration, which has been shown to be less susceptible to fouling and is
consequently favoured in applications using fouling-prone feed solutions [52]. Using different
membrane surfaces will provide insight in foulant-membrane interactions governing fouling as well.
In addition, applying the high ionic strength of seawater will enable an assessment of the fouling
potential of state-of-the-art membranes in FO seawater desalination applications.
3.1. Single and combined fouling
3.1.1. Comparison CTA, TFC, and POR In this study, three different membranes were compared: a cellulose tri-acetate (CTA) membrane, a
thin film composite (TFC) membrane (both from HTI), and a novel membrane obtained under a
confidentially agreement (referred to as POR). To ensure proper comparison between the fouling
propensity of the different membranes, it is essential that they are compared under similar operating
conditions. Therefore, the draw solution concentration was adjusted to obtain the same initial flux in
each set of experiments, which was 16.84 ± 0.71 L/m²/h (Table 8). Due to the variation in required
bulk osmotic pressure difference between the membranes (mainly caused by the difference in ICP,
which is indicated by the S-values) to obtain the similar initial fluxes (for example, for CTA the
required draw concentration to reach the flux of 16.84 L/m²/h is 3.7 M, whereas for the POR
membrane the draw concentration was only 1.5 M), the osmotic driving force will drop in different
ways for the different membranes (i.e. more easily in case of POR). Thus, to minimise this effect, only
the first 10 h of the fouling experiments were compared since the driving force is still more or less
similar for all membranes in this short time interval.
Figures 17a, b, and d represent the fouling trends of the different membranes for the different
foulants (alginate, silica colloids and their combination). The CTA, TFC, and POR membranes are all
compared in AL-FS mode, but for comparison, also the flux decline for the CTA membrane in AL-DS
mode is shown. The CTA membrane is the only membrane that is not a thin-film composite, and that
can thus be used in both membrane orientations. In addition to Figure 17, the final flux decline
percentages (compared to the baseline) are displayed in Table 8 for further comparison. The final flux
decline percentages in case of fouling with silica colloids showed the following trend: POR (24%) >
50
TFC (3%) > CTA-FO (0%), indicating a severe loss is permeate flux when the POR membrane was
applied, while this was clearly not the case when applying the CTA or TFC membrane. Alginate fouling
resulted in a clear distinction between the POR membrane and the two other membranes as well. A
final flux decline of 30% was noted for POR, whereas only 5% and 7% for TFC and CTA, respectively.
When both alginate and silica colloids were combined, the flux decline trend in case of CTA and TFC
again decreased in a much lower extent (13% and 11%, respectively) compared to the flux decline
trend of POR (34%). In all cases alginate, silica colloids, and combined fouling exhibited rather low
initial flux decline rates for CTA and TFC, whereas in case of POR higher initial flux decline rates could
be observed (Table 9). Furthermore, the flux declined to a higher extent when applying the
combination of the foulants. However no clear evidence of synergy between both alginate and silica
colloids could be observed, similar to the observation made in Paragraph 2.1.
Table 7. Respective draw concentrations and final flux decline percentages (compared to the baseline after 10 h) for each membrane (in AL-FS mode). CTA-PRO refers to the case were the CTA membrane was applied in AL-DS mode, while CTA-FO indicates the membrane was applied in AL-FS mode
Membrane Draw concentration (M) Final flux decline (%)
ALG ST-ZL Comb
CTA-PRO 3.5 32 15 39
CTA-FO 3.7 7 0 13
TFC 4.2 5 3 11
POR 1.5 30 24 34
It is clear that the POR membrane shows a higher fouling tendency in case of either alginate, silica
colloids, and their combination compared the CTA and TFC membrane. In addition, no substantial
differences occurred between the CTA and TFC membrane when examining the flux decline trend of
the different foulants (although the CTA membrane was clearly not fouled by the silica colloids at all).
These results indicate that even though the POR membrane outperforms HTI’s membranes (CTA and
TFC) in terms of performance parameters (Table 3), the opposite is true under fouling conditions as
more rapid flux decline rates and higher losses of flux were observed for each foulant in case of POR.
Most likely, this is due to modifications of the active layer in terms of surface functionality and/or
surface roughness compared to the CTA and TFC membranes, as these two factors have been
considered of great importance (regarding fouling propensity) in literature [70, 72]. At this point, it is
difficult to determine which of both factors seems more likely. Further on in this study (Paragraph
3.2.), the XDLVO approach was applied in an attempt to clarify this matter.
51
Table 8. Overview initial flux decline rates for each foulant-membrane combination.
Membrane Initial flux decline rate (h-1)
ALG ST-ZL Comb
CTA 0.0079 0 0.0137
TFC 0.0098 0.0112 0.0164
POR 0.0297 0.0193 0.0304
Time (h)
0 2 4 6 8 10
No
rmal
ised
flu
x
0,5
0,6
0,7
0,8
0,9
1,0
1,1
ALG
ST-ZL
Comb
Baseline
Time (h)
0 2 4 6 8 10
No
rmal
ised
flu
x
0,5
0,6
0,7
0,8
0,9
1,0
1,1
ALG
ST-ZL
Comb
Baseline
Time (h)
0 2 4 6 8 10
No
rmal
ised
flu
x
0,5
0,6
0,7
0,8
0,9
1,0
1,1
ALG
ST-ZL
Comb
Baseline
Time (h)
0 2 4 6 8 10
No
rmal
ised
flu
x
0,5
0,6
0,7
0,8
0,9
1,0
1,1
ALG
ST-ZL
Comb
Baseline
A B
C D
Figure 15. Overview permeate flux decline curves in case of fouling a (a) CTA (AL-FS mode), (b) TFC, (c) CTA (AL-DS), and (d) POR membrane with single foulants (1 g/L ST-ZL silica colloids and 200 mg/L alginate) and their combination in the presence of mono and divalent salts (0.476 M NaCl and 0.008 M CaCl2). Each fouling experiment was performed in AL-FS mode (except the CTA membrane depicted in (c)) at a cross-flow of 30 l/h for 24 h.
3.1.2. Comparison AL-DS and AL-FS mode Since the CTA membrane was subjected to the exact same fouling conditions in both operational
modes (AL-DS and AL-FS) and the initial fluxes were in the same order of magnitude, a fair
comparison could be made of the fouling propensity of both orientations. For clarity, only the first 10
h of Figure 12 were redisplayed in Figure 17c. It is clear that the same order in the magnitude of the
52
flux decline between the foulants occurred (which was already stated in the previous paragraph),
although a much higher flux decline occurred when the CTA membrane was positioned in the AL-DS
mode (for each foulant). Moreover, the silica colloids did not foul the CTA membrane in the AL-FS
mode, whereas a substantial flux loss of 15% was observed in the AL-DS mode after 10 h. The
existence of such a major difference between both configurations is most likely due to the porous
layer of the support which is facing the foulants in the AL-DS configuration, and provides a more
rough and to some extent penetrable surface for the foulants compared to the smooth and dense
active layer [52]. Consequently, foulants are more readily trapped by the membrane structure in the
AL-DS mode, which clearly has an important effect on colloidal fouling with silica, but also enhances
alginate and combined fouling. In case of individual alginate fouling, this is in good agreement with
literature [52, 53, 64]. No reports were found where silica colloids and combined fouling with the
colloids and alginate were applied on side of the porous layer.
3.2. Relation between interfacial free energies of adhesion and
membrane fouling In Paragraph 2.3., the relation between the interfacial free energies (of adhesion and cohesion) and
the fouling propensity of the CTA membrane in AL-DS mode was established for the different
foulants. Here, a similar analysis is made for the CTA, TFC, and POR membrane (all in AL-FS mode).
The respective surface tension components and the interfacial free energies for the different
membranes and different foulants, as determined from contact angle measurements, are listed in
Tables 6 and 7. For each membrane, similar to the observed adhesion energies between the alginate
and the CTA membrane in AL-DS mode, clear negative values were obtained for the interfacial
energies between the alginate and the membranes – indicating adhesive forces; while strongly
positive values were acquired for the colloids – indicating repulsion. Again, the interfacial energy
between the combination of the foulants and the membrane is situated in between these extremes
(Table 7). Nevertheless, substantial differences in the extent of adhesion energies occured,
depending on the type of membrane.
In case of silica colloids, the CTA membrane (when operated in AL-FS mode) exhibited highly positive
adhesion energies (28.68 mJ/m²), which is in contrast with the only slightly positive value for the POR
membrane (6.61 mJ/m²). This suggests that strong repulsive forces exist between the active layer of
the CTA and the silica, while much less repulsion is be expected in case of POR. This clearly
corresponds to the fouling patterns displayed in Figures 17a and d, where no colloidal fouling
occurred for the CTA in AL-FS mode, and a tremendous flux decline appeared for POR. The TFC
53
membrane showed an intermediate trend for both the adhesion energies (18.02 mJ/m²) and the
resulting fouling behaviour (Figure 17b).
Only slight dissimilarities can be noticed when comparing the adhesive energies obtained for alginate
and the different membranes. For CTA, TFC, and POR, the free energies of adhesion were -9.42
mJ/m², -7.32 mJ/m², and -9.69 mJ/m², respectively. The clearly negative energies imply that alginate
fouling is able to occur spontaneously no matter which membrane is applied. This is in accordance
with the observed flux decline trends, where alginate fouling took place regardless of the membrane
surface. However, the distinct difference in fouling propensity between the POR membrane and
those of HTI, was less clear in case of the predicted negative adhesion energies
Also for combined fouling, the trend depicted by the adhesive forces between the respective
membranes and the combination of both foulants is in accordance with the observations. Both CTA
and TFC depicted (strongly) positive free energies of adhesion (12.69 mJ/m² and 6.80 mJ/m²,
respectively), while the POR membrane showed a much lower and slightly negative value (-0.40
mJ/m²). Consequently, a larger permeate flux decline could be expected for the POR membrane,
which is confirmed by the obtained flux decline trends.
In all cases, the differences (in terms of the extent of the adhesion energies) depicted by the results
of the XDLVO calculations largely matched with the observed dissimilarities (in terms of the losses in
permeate flux decline) for each foulant-membrane combination. This confirms the usefulness of the
interfacial energies to explain and support differences in fouling trends. In addition, the adhesion
energies strongly suggest that the distinctly larger fouling propensity of the POR membrane
compared to the CTA and TFC membrane is mainly due differences in surface functionalities rather
than surface roughness (which was not clear in Paragraph 3.1.1.). However, the roughness has to be
specifically examined to be conclusive about this matter, which requires further research (for
example through AFM analysis).
3.3. Membrane applicability in stand-alone FO desalination The performance of the applied membranes can be assessed relative to each other in terms of
intrinsic parameters (A, B, S, and R )and overall fouling propensity. Both criteria have a direct effect
on the permeate flux and consequently on the overall yield of the stand-alone FO desalination
system (Part II, Paragraph 1.1.2.). Table 10 provides an overview of how each membrane fulfils these
criteria (on a scale relative to each other), thereby combining the results of Paragraphs 1.1. and
3.1.1. The evaluation of the intrinsic parameters is based primarily on the structure factor, as this
value influences the permeate water flux to the largest extent [105]. The TFC membrane scores the
54
worst on this criterion, with the largest S factor, followed by CTA and POR, respectively. Since POR
outperforms both TFC and CTA, which was also observed by examining the initial draw concentration
(Paragraph 3.1.1.), the highest score was given regarding the intrinsic parameters (Table 10),
indicating the contrast. The assignment of scores in terms of fouling propensity was based on the
clear dissimilarities outlined in Paragraph 3.1.1, where POR was found to be highly prone to fouling
compared to CTA and TFC.
Examination of the performance criteria allows to state that both the novel TFC and POR membranes
are currently not yet suitable for stand-alone FO desalination as a result of contrasting reasons.
Despite the potentially high permeate fluxes that can be obtained in case of clean POR membranes,
drastic losses in performance will occur in practice due to the elevated fouling propensity (which will
be even more pronounced at the higher fluxes that the POR is able to reach). In terms of overall
performance, the TFC stands in total contrast with POR, since the high levels of ICP will limit the
maximum achievable flux for the clean membrane in the stand-alone process to a large extent,
whereas membrane fouling appeared to be the much less of a problem. Within these extremes, it is
reasonable to state that the overall performance of the traditional CTA membrane is still good, even
though for a given draw solution concentration, the product yield of the clean membrane will be
substantially lower than that of the clean POR membrane (but still higher than TFC, which is
remarkable). Thus, unless changes can be made in the current properties of the CTA, TFC, and POR
membranes, it appeared that no membrane is yet entirely suited for the purpose of seawater
desalination by FO. For CTA and TFC, this would imply reducing the thickness of the porous layer,
without changing the properties of the active layer (in order to sustain its relatively low fouling
propensity). Reports can be found in literature which state that a significantly lower structure factor
was obtained compared to the commercial CTA and TFC membranes by optimising the
manufacturing process [49, 87, 93], indicating there is still room for improvement for these
membranes. In case of the POR membrane, further research will have to focus on increasing its
fouling resilience in order to improve its applicability (unless a pre-treatment step for the feed
stream would be installed, which would of course increase the cost of the operation).
Table 9. Assessment of membrane performance (in AL-FS mode) based on two criteria: (1) intrinsic parameters (A, B, S, and R) and (2) fouling propensity. (1) is rated as bad, average or good, while (2) is rated as low, medium or high.
Membrane Performance
Intrinsic parameters Fouling propensity
CTA bad - average low
TFC bad low POR good high
55
Part V: Conclusions and future prospects
The first goal of this work was to determine the factors that govern combined organic and colloidal
fouling of forward osmosis membranes (with alginate and silica colloids as the respective model
foulants) at the ionic strength of seawater (represented by 0.0476 M NaCl and 0.008 M CaCl2,
resulting in a total ionic strength of 0.5 M). It was shown that regardless of the applied foulant (either
alginate, silica colloids, or their combination), no substantial losses in permeate flux – and thus no
substantial fouling – occurred in the absence of mono and divalent salts. However, increasing the
ionic strength to 0.5 M resulted in significant fouling behaviour, and distinct differences between the
foulants.
- Calcium was considered to be the main contributor to these observations, causing severe
alginate fouling due to intermolecular bridging and to a lower extent colloidal fouling of silica
due to colloid destabilisation.
- Remarkably, the combination of both foulants (alginate and silica colloids) resulted in a flux
decline trend similar to that of individual alginate fouling, even though the silica colloids
were in abundance compared to alginate (1 g/L versus 0.2 g/L).
- Results from consecutive fouling experiments suggested that the hydraulic resistance of the
combined fouling layer was the main cause for the observed flux decline.
- In addition, alginate aggregates seemed to adsorb on the silica colloids, which could imply
that the colloids became embedded within the alginate layer during combined fouling. The
adsorption is probably also the reason why alginate could shear away colloids from the
membrane surface during the consecutive fouling experiments, thus removing the colloidal
fouling.
From the observations above, it is believed that combined fouling with alginate and silica colloids at
the given (high) ionic strength is dominated by alginate-calcium complexation, possibly
overwhelming the presence of the silica colloids.
In the second phase of this research, a comparison was made between the performance of a
standard CTA membrane and novel TFC and POR membranes in terms of fouling propensity (again
using alginate, silica colloids, and their combination as foulants). Prior to the fouling experiments, the
intrinsic parameters were determined for each membrane. From these results, it was observed that
the (clean) POR membrane exhibited superior water permeability (i.e. a high A-value) and a drastic
reduction of ICP (i.e. a low S-value) compared to both CTA and TFC. The fouling experiments, on the
other hand, showed that the POR membrane exhibited the least fouling resilience, whereas the TFC
56
and CTA membrane were substantially less prone to fouling. Alginate for example, caused a flux
decline up of 30% in case of the POR membrane, which is in sheer contrast with the 7% and 5% flux
declines noted for CTA and TFC, respectively. These results indicate that even though the POR
membrane outperformed the CTA and TFC membranes in terms of membrane characteristics
(primarily by reducing the performance-limiting effect of ICP), the opposite is true under fouling
conditions. When the performance of the different membranes was assessed in terms of both
membrane properties and fouling propensity (relatively to each other), it was considered that
currently, none of the applied membranes can be deemed (entirely) suitable for a stand-alone FO
desalination process on seawater, as either ICP or fouling drastically reduce the permeate flux.
Throughout this work, the XDLVO approach was used to further investigate fouling mechanisms and
to investigate the link between flux decline rates and foulant-membrane and foulant-foulant
interactions. These interactions were determined by calculating the free energies of adhesion and
free energies of cohesion from contact angle measurements. It was observed that clear correlations
could be found between the initial flux decline rates and adhesive energies on one hand, and
between the later flux decline rates and cohesive energies on the other hand. The results from the
interfacial energy calculations thus confirmed the higher experimentally observed flux declines in
case of both alginate and combined fouling compared the silica colloids. When the different
membranes were compared, it was shown that the differences in the extent of the adhesion energies
largely matched with the observed experimental dissimilarities (in terms of losses in permeate flux
decline) for each foulant-membrane combination. In addition, the adhesion energies strongly suggest
that the distinctly larger fouling propensity of the POR membrane compared to the CTA and TFC
membrane was mainly due differences in surface functionalities (rather than surface roughness).
Overall, these results confirm the usefulness of the interfacial energies to explain and support
differences in fouling behaviour, and to shed more light on fouling mechanisms.
Future research regarding fouling behaviour and/or mechanisms in FO should focus more on
combined fouling, in order to better comprehend fouling behaviour in real-life applications. Since
membrane fouling can be dominated by one particular foulant or synergy between foulants can exist,
a distinct dissimilarity can occur between individual fouling and combined fouling behaviour. In
addition, such research should expand further on methodologies such as the XDLVO approach, which
can support/explain the observed fouling trends/propensities by determining the underlying
interactions between foulant and membrane, as well as between foulants. Furthermore, research
regarding membrane development should incorporate the analysis of a membrane’s fouling
propensity next to improving its intrinsic membrane properties. In that case, a more complete (and
more practical) assessment can be made regarding the applicability of a newly developed membrane
57
for practical applications. In particularly, the very promising POR membranes should be adapted to
reduce its fouling propensity, without reducing its outstanding membrane properties.
58
59
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Appendices
APPENDIX A. CONTACT ANGLE MEASUREMENTS Contact angles are measured using a contact angle goniometer, which was a Krüss DSA 10-Mk2 in
this study. In Figure 18, the main parts of this device are highlighted. The light bulb ensures a good
level of contrast, a clear images emerges on the screen, as the camera is connected to a computer.
This visualisation enables the measurement of the contact angles using the Drop Shape Analysis
program (version 1.8).
Figure 16. Picture of the Krüss DSA 10-Mk2 contact angle measurement device with a camera (1), sample platform (+ sample) (2), needle (for droplet deposition) (3), and light source (4).
The exact angle is measured between the baseline and the tangent through the point where the
droplet touches the surface, which is illustrated by Figure 19. The height and inclination of baseline
can be manually adjusted to correct for small irregularities in the membrane/foulant surface.
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Figure 17. Illustration of the contact angle calculation principle. The angle is measured between the baseline (blue) and the tangent through the point where the droplet touches the surface (red).
69
APPENDIX B. DETERMINATION OF A, B, R, AND S In order to calculate A, a 6 L solution of 0.01 M NaCl was prepared in DI water after which it was
circulated on the active layer side (feed side) of the membrane at 60 L/h. As no natural driving force
is present, a hydraulic pressure of 2, 4, and 6 bar was applied on the feed side to establish
permeation across the membrane. Consequently, A can be calculated by Eq. 1 (where disappears
because not osmostic driving force exists), as Jw is found by Eq. 11.
R was calculated using Eq. 12 in which the concentration of salt at both sides of the membrane could
be found by measuring the conductivity of the solution at each side (it is known that the salt
concentration of a solution in linearly related to its conductivity up to 1 M).
B was calculated using Eq. 11 , which required values for the parameters Jw, R, and k. In case of Jw and
R, the abovementioned procedure was followed. To find k, the procedure presented in Part II,
Paragraph 2.1., is required.
Finally, S was found by Eq. 14 by implementing the necessary constants.
During each test, a RO feed spacer was used at both sides of the membrane.
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APPENDIX C. SEM IMAGES OF CLEAN MEMBRANES To assess the structure and morphology of the membranes and the different fouling layers, a
visualization technique was applied, using a Jeol JMC-5000 scanning electron microscope (Nikon
Instruments Inc., U.S.A.). Each membrane sample was dried at 60°C for at least 30 min to remove
moist. Subsequently, the samples were coated with gold to provide electrical conductivity and
prevent charging during imaging, which improves the clarity of the images.
The active layer, porous supper layer, and the cross section of a CTA membrane are presented in
Figures 20a, b, and c respectively. It can be observed that the active layer is dense compared to the
loose porous support layer, which is typical for an assymetrical membrane. The woven PET fabric is
embedded in between these layers, which is shown in Figure 20c.
Figure 18. SEM images from the active layer (a), porous layer(b), and cross section (c) of a CTA membrane.
A B
C
72
In case of the TFC membrane (Figure 21), again the fabric (supposedly also PET) loose porous support
layer is embedded in the loose support layer. Although the porous layer is much looser than the
porous layer of the CTA membrane (which claimed by the manufacturer and observed during fouling
experiments), the image is not clear enough the indicate this.
Figure 19. SEM images from the active layer (a), porous layer(b), and cross section (c) of a TFC membrane.
A B
C
A
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APPENDIX D. (SEM) IMAGES OF FOULED MEMBRANES The appearance of foulants on the membrane surface was noticed after every fouling run, regardless
of the foulant solution applied. A few images of those fouling layer are given below (Figure 22).
Figure 20. Images of fouled (a) POR, (b) CTA, and (c) TFC membranes (in AL-FS mode) when both alginate and silica colloids were applied in the highly saline feed solution.
The samples prepared for contact angle analysis of the combination of both alginate and silica
colloids are shown below. A top layer of silica colloids can be observed in every case, which led the
remarked underestimation of the adhesion energies (Part IV, Paragraph 2.3.2.).
Figure 21. Images of samples prepared for contact angle analysis of combined fouling with alginate and silica colloids.
A B
C
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Figure 24 displays SEM images of different the fouling layers in case of fouling with alginate (a), silica
colloids (b), and their combination (c).
Figure 22. Top view SEM images of CTA membrane fouled with alginate (a), silica colloids (b), and their combination (c). In all cases, the feed solutions contained 0.476 M NaCl and 0.008 M CaCl2.
A B
C
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APPENDIX D. DETERMINITION OF ZETA POTENTIAL AND
HYDRODYNAMIC PARTICLE SIZE
D1. Particle sizing The mean hydrodynamic diameter of the alginate macromolecules, silica colloids, and the combined
aggregates in the different feed solutions was determined by dynamic light scattering (DLS) using a
Malvern photon correlation spectrometer (Malvern Instruments, UK). The concentration of alginate
(0.5 g/L) used in the size determination was higher than that used in the fouling experiments to
enhance sample detection, since 200 mg/L was too dilute. It is worth mentioning that this
concentration can influence the observed changes in alginate particle size due to multiple lights
scattering resulting in over-estimated particle size. Therefore, to determine this effect, the alginate
particle size was measured at 0.3 g/L, 0.75 g/L and 1 g/L and the resulting values were not
significantly different from each other and the chosen 0.5 g/L that was used for this work. Thus we
assumed that the size and zeta potential trends observed could be extrapolated down to 0.2 g/L.
D2. Zeta potential measurements The alginate, silica, and combined aggregates surface charges with and without the presence of
mono and divalent salts were determined from electrophoretic mobility, measured using a Malvern
Zetasizer300 HS series.(Malvern Instruments, UK).
D3. Results Table 10. Measured zeta potential and hydrodynamic diameter of alginate and silica colloids in different solutions.
Solution/Membrane Zeta potential Hydrodynamic diameter
mV nm
ALG -47.5 66.02
ST-ZL -54.4 139.9
ALG + ST-ZL -72.15 184.9
ALG (0.5 M)* -8.85 535.4
ST-ZL (0.5 M)* -14.2 156.35
ALG + ST-ZL (0.5 M)* -15.2 177.8
CTA-PRO (0.1 M)** -10.78*** N.A.
CTA-FO (0.1 M)** -3.80*** N.A.
CTA-PRO (0.5 M)* 19.83*** N.A.
CTA-FO (0.5 M)* 6.68*** N.A.
* The total ionic strength is 0.5 M, composed out of 0.476 M NaCl and 0.008 M Ca Cl2. ** The total ionic strength is 0.1 M (NaCl). *** Results obtained from Motsa, et al. [64].
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