DESALINATION BY MEMBRANE DISTILLATION: FABRICATION OF … · Fig.4.4 Schematic diagram showing the...
Transcript of DESALINATION BY MEMBRANE DISTILLATION: FABRICATION OF … · Fig.4.4 Schematic diagram showing the...
DESALINATION BY MEMBRANE DISTILLATION:
FABRICATION OF HIGH PERFORMANCE MEMBRANES
SINA BONYADI
(B. Eng. (Chemical) (Hons.), Amirkabir University of Technology)
A THESIS SUBMITED FOR
THE DEGREE OF MASTER OF ENGINEERING
DEPARTMENT OF CHEMICAL & BIOMOLECUALR
ENGINEERING
NATIONAL UNIVERSITY OF SINGAPORE
2008
Acknowledgement
First of all, I would like to express my deepest heartfelt appreciation to my supervisor Prof. Tai-
Shung Chung Neal, in Dept. of Chemical & Biomolecular Engineering of NUS, for his excellent
guidance, enthusiastic encouragements and invaluable suggestions throughout my two year
master study. From him, I have learnt a great deal on both research knowledge and active work
spirits.
I am especially grateful to Prof. William B. Krantz Isac mayor professor in Dept. of Chemical &
Biomolecular Engineering of NUS for his wise guidance regarding my first paper in the journal
of membrane science.
Special thanks are also due to Mr. NG Kim Po the workshop chief at NUS Chemical Engineering
department for the fabrication of membrane distillation experimental set up.
I also want to take this opportunity to give my sincere thanks to all the colleagues in my research
group for their kind assistance. My association with them is a memorable part of my experience
and study at NUS.
I gratefully acknowledge A*STAR for providing me an opportunity to pursue my Master degree
and research scholarship.
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Table of Contents
Acknowledgement………………………………………………………………………………....i
Summary…………………………………………………………………………………………..vi
List of Tables…………………………………………………………………………………….viii
List of Figures…………………………......……………………………………………………...ix
List of symbols…………………………………………………………………………………...xii
Chapter 1. Background Review and Objectives
1.1 Introduction…………………………………………………………….………………...1
1.2 Desalination Processes………………………………………………….………………..3
1.3 Alternative desalination processes………………………………….................................6
1.4 Introduction to membrane distillation as an alternative desalination approach…….........7
1.5 Membrane distillation configurations……………………………………………………8
1.6 Temperature polarization phenomenon…………………….............................................9
1.7 Wetting phenomenon………………………………………………………………….....9
1.8 Applications of membrane distillation……………………………………………….....10
1.9 Membrane distillation advantages and drawbacks…………………………………..….10
Chapter 2. Literature Review
2.1 Overview on MD literature……………………………………………………………...12
2.2 Literature review on membrane fabrication for MD…………………………………….13
2.3 Research objectives……………………………………………………………………...17
Chapter 3. Theory and model development
3.1 Mass transfer……………………………………………………………………………..19
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3.2 Heat transfer……………………………………………………………………………...22
3.3 Characteristics of a high performance MD membrane…………………………………....24
3.3.1 High membrane permeability……………………………………………………..24
3.3.2 High membrane wetting resistance and long term stability……………………....26
3.3.3 Suitable membrane geometry and dimensions……………………………………26
3.3.4 Hydrophilic layer porosity and thermal conductivity…………………………….27
Chapter 4. Experimental
4.1 Materials………………………………………………………………………………...28
4.2 Dope preparation………………………………………………………………………...29
4.3 Fabrication of flat sheet membranes…………………………………….........................30
4.4 Fiber spinning……………………………………………………………………………30
4.5 Morphology study of hollow fibers by SEM…………………………….........................31
4.6 Contact angle measurements……………………………………………..........................31
4.7 Porosity measurement……………………………………………………………………32
4.8 Pore-size distribution measurement……………………………………………………...33
4.9 Gas permeation test………………………………………………………………………33
4.10 Polymer flow observation by a high magnification camera……………………………..34
4.11 Video microscopy flow visualization……………………………………........................34
4.12 Module Fabrication………………………………………………………........................34
4.13 DCMD experiments……………………………………………………………………...35
Chapter 5. Fabrication of Dual Layer Hydrophilic-Hydrophobic Hollow Fibers
5.1 Effect of coagulant on the surface morphology of PVDF membranes…………………..37
5.2 First batch of fiber spinning……………………………………………………………...39
5.2.1 Membrane morphology…………………………………………………………...40
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5.2.2 DCMD performance……………………………………………………………...41
5.3 Second batch of hollow fiber spinning…………………………………………………..41
5.3.1 Hollow fibers morphology ………………………………………........................43
5.3.2 Pore-size distribution…………………………………………….........................45
5.3.3 Gas permeation and porosity measurement tests…………………........................45
5.3.4 Contact angle measurements……………………………………………………...45
5.3.5 DCMD results…………………………………………………………………….46
5.4 Summary………………………………………………………………………………...50
Chapter 6. Investigation of Corrugation Phenomenon in the Inner Contour of Hollow
Fibers during the Non-solvent Induced Phase-Separation Process
6.1 Introduction………………………………………………………………………………51
6.2 Experimental observations……………………………………………….........................53
6.3 Discussion………………………………………………………………………………..58
6.3.1 System description…………………………………………………………………...59
6.3.1.1 Phases I1, I2 or O1, O2…………………………………….……………………60
6.3.1.2. Phase I3 or O3…………………………………………………………………60
6.4 Possible instability mechanisms………………………………………………………...61
6.4.1 Hypothesis 1. (Mass transfer and hydrodynamic instability)………………………..62
6.4.2 Hypothesis 2. (Elastic and Buckling Instability)………………….............................64
6.5 Effect of air-gap distance………………………………………………………………...66
6.6 Effect of bore fluid composition…………………………………….…………………...68
6.7 Effect of external coagulant………………………………………….…………………..68
6.8 Effect of take-up speed………………………………………………..............................68
6.9 Effect of dope concentration…………………………………………..............................69
6.10 Summary………………………………………………………………………………...69
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Chapter 7. Conclusion…………………………………………………………………………..70
Bibliography……………………………………………………………………………………..72
Appendix…………………………………………………………………………………………84
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Summary
For the first time, co-extrusion was applied for the fabrication of dual layer hydrophilic-
hydrophobic hollow fibers especially for the direct contact membrane distillation (DCMD)
process. The effect of different non-solvents on the morphology of the PVDF membranes was
investigated and it was found that weak coagulants such as water/methanol (20/80 wt%) can
induce a 3-dimensional porous structure on PVDF membranes with high surface and bulk
porosities, big pore size, sharp pore size distribution, high surface contact angle and high
permeability but rather weak mechanical properties. Hydrophobic and hydrophilic clay particles
were incorporated into the outer and inner layer dope solutions, respectively, in order to enhance
mechanical properties and modify the surface tension properties in the membrane inner and outer
layers. Different membrane characterizations such as pore size distribution, gas permeation test,
porosity and contact angle measurements were carried out as well. Ultimately, the fabricated
hollow fibers were tested for the DCMD process and flux as high as 55 kg/m2hr and energy
efficiency of 83% at 90 °C was achieved in the test. The obtained flux is much higher than most
of the previous reports, indicating that the application of dual layer hydrophilic-hydrophobic
hollow fibers may be a promising approach for MD.
In the second part of this research, by proposing a novel mechanism, we revealed one of the most
controversial issues in the hollow fiber fabrication process regarding the instability leading to the
deformed cross-section of fibers fabricated through nonsolvent induced phase separation. We
analyzed possible instability mechanisms based on our experimental observations and then
postulated that the principal instability occurs in the external coagulation bath where the rigid
precipitated polymer shell in the dope and bore fluid interface is buckled due to a generated
pressure. The pressure is postulated to be induced in the nascent fiber outer layer as a result of
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diffusion/convection, precipitation, densification and shrinkage In addition, the effect of some
spinning conditions such as air-gap distance, bore fluid composition, take-up speed, external
coagulant and dope concentration on the final shape of the fiber cross-section have been
investigated. The proposed mechanism was in good qualitative agreement with all our
observations.
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List of Tables
Table.2.1 Summary of commercial membranes applied by some studies in the literature
Table.4.1 Specifications of Cloisite particles
Table.5.1 Total solubility parameter (δt) of water, methanol, NMP and PVDF
Table.5.2 Spinning conditions applied for the first batch of spinning
Table.5.3 DCMD operating conditions and the obtained flux for the fabricated fibers
Table.5.4. Comparison of the maximum flux obtained in this study with the literature for DCMD processes with a hollow fiber configuration
Table.6.1 Spinning conditions of hollow fiber membrane fabrication
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List of Figures
Fig.1.1 Distribution of Earth’s water
Fig.1.2 Fresh water resources
Fig.1.3 Schematic presentation of a Multi-Stage Flash desalination plant
Fig.1.4 Schematic diagram of world’s desalination plants capacity percentage by 1998
Fig.1.5 Typical Costs for a Reverse-Osmosis Desalination Plant
Fig.1.6 Schematic diagram representing the separation mechanism involved in MD
Fig.1.7 Schematic diagrams representing different configurations of the MD process
Fig.1.8 Schematic diagram of temperature polarization phenomenon in DCMD
Fig.2.1 Schematic diagram representing a typical melt spinning process
Fig.2.2 A composite hydrophilic-hydrophobic membrane before the MD test (left) A composite hydrophilic-hydrophobic membrane during the MD test (right)
Fig.2.3 Schematic picture of a dual layer hydrophilic-hydrophobic fiber
Fig.3.1 Schematic DCMD process with dual layer hydrophilic-hydrophobic hollow fibers
Fig.4.1 Chemical structure of PVDF polymer
Fig.4.2 Chemical structure of PAN polymer
Fig. 4.3 Schematic representation of the fiber spinning line
Fig.4.4 Schematic diagram showing the steps involved in fabrication of lab scale membrane modules
Fig.4.5 Schematic experimental set-up applied for direct contact membrane distillation process
Fig.5.1 SEM pictures from the top surface (facing the coagulant) of the PVDF flat sheet membranes for different coagulant compositions
Fig.5.2 SEM pictures representing the structure of first spinning batch fibers
Fig.5.3 Delamination phenomenon during the DCMD process using the first batch of fibers
Fig.5.4 The spinning conditions applied in the second batch of spinning process
Fig.5.5 Cross section morphology of fibers fabricated through the second batch of spinning
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Fig.5.6 SEM micrographs showing the surface morphology of the dual layer fibers
Fig.5.7 Pore size distribution of the fabricated fibers
Fig.5.8 Contact angle measurements of the fabricated fibers
Fig.5.9 Flux and energy efficiency obtained in the DCMD process using the fabricated fibers and flux comparison with the literature data using hollow fiber membranes
Fig.6.1 Irregular shape in the cross section of the PAN 17% (left) and PVDF 20% (right) hollow fibers fabricated through wet spinning, Bore fluid composition of 40% NMP/water (NMP wt.%), free fall take up rate Fig.6.2 Cross section of spun fibers from 17 wt.% PAN solution, 40 wt.% NMP Bore fluid with different air-gap distances
Fig.6.3 Cross section of wet spun fibers from 17 wt.% PAN solution with different solvent amount in the bore fluid
Fig.6.4 Cross section of wet spun fibers from 17 wt.% PAN solution and 40 wt.% NMP bore fluid using different external coagulants: (A) IPA (B) water
Fig.6.5 Cross section of spun fibers with different take-up rates from 17 wt.% PAN solution, 40 wt.% NMP bore fluid, 7 cm air-gap
Fig.6.6 Cross section of spun fibers from different concentrations of PAN solution, 40 wt.% NMP aqueous mixture as bore fluid, with 3.5 cm air-gap and free fall (A) 13 wt.% PAN solution (B) 17 wt.% PAN solution (C) 22 wt.% PAN solution
Fig.6.7 Cross section of spun fibers from 20 wt.% PVDF solution, 40 wt.% NMP Bore fluid with different air-gap distances
Fig.6.8 Cross section of wet as-spun hollow fibers from 17 wt.% PAN solution and bore fluid containing 40 wt.% NMP
Fig.6.9 Magnified die swell pictures of spinning with 17 wt.% PAN solution, Bore fluid containing 40 wt.% NMP and 7 cm air-gap distance, Different bore flow rates
Fig.6.10 (A) Penetration of coagulant into the casting solution after 0.2 second using flow visualization photos taken on flat membrane inversion as an example, (B) Schematic regions in the extruded nascent fiber
Fig.6.11 Instability associated with proposed hydrodynamic and mass transfer mechanism
Fig.6.12 Schematic picture showing the exertion of inward radial forces generated by the shrinkage of nascent fiber outer layer in the external coagulation bath
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Fig.6.13 Close agreement between the spun fibers cross sectional geometry (A) and the predicted postbuckling shapes of a long elastic cylindrical shell (B)
Fig.6.14 Energy as a function of the dimensionless group including the external pressure for various modes of deformation
Fig.6.15 Equilibrium state transitions of the buckled elastic shell as a function of spinning conditions
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List of Symbols
A, B, C Antoine Equation Constants
Afo hollow fiber outer surface area (m2)
Afi hollow fiber inner surface area (m2)
Ce Membrane lumped diffusive coefficient (m2/s)
Cf feed heat capacity (kJ/kg ºK)
Cp permeate heat capacity (kJ/kg ºK)
Cv water vapor heat capacity at the constant volume (kJ/kg ºK)
Dwk water vapor Knudsen diffusion coefficient (m2/s)
Dw-a water vapor-air molecular diffusion coefficient (m2/s)
di fiber inner diameter (m)
F flexural rigidity
G Young’s modulus of elasticity (Pa)
Hp permeate side overall heat transfer coefficient (J/m2°C)
h Shell thickness (m)
hf feed-side heat transfer coefficient (J/m2°C)
hp permeate-side heat transfer coefficient (J/m2°C)
kair thermal conductivity of air (m2/s)
ki thermal conductivity of the fiber inner layer (m2/s)
k thermal conductivity of the fiber outer layer (m2/s)
kB Boltzmann constant
km thermal conductivity of the membrane matrix (m2/s)
kw water thermal conductivity (m2/s)
L fiber length (m)
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Ma air molecular weight
Mw water molecular weight
fm feed mass flow rate (kg/hr)
pm permeate mass flow rate (kg/hr)
Nw water vapor mass flux (kg/m2 hr)
Nu Nusselt number
n number of hollow fibers in the membrane module
P Dynamic pressure (Pa)
Pa air partial pressure (Pa)
Pw water vapor partial pressure (Pa)
*P Critical buckling pressure (Pa)
Pr Prandtl number
q Buckling mode
r radial coordinate (m)
rp pore radius
R universal gas constant
Rh cylinder radius (m)
Ri fiber inner radius (m)
Rm hydrophilic layer outer radius (m)
Ro fiber outer radius (m)
Re Reynolds number
Red Reynolds number based on fiber inner diameter as the characteristic length
T temperature (C)
Tfi feed inlet temperature (C)
Tpi permeate inlet temperature (C)
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t Time (s)
U Potential Energy (kg m/s2)
u
Velocity vector (m/s)
V fiber empty volume (m3)
v Poisson’s ratio
w1 weight of the sample before immersion in kerosene (kg)
w2 weight of the sample after immersion in kerosene (kg)
xw mole fraction of water in the feed
z axial coordinate (m)
σ collision diameter of the molecule (m)
δm membrane thickness (m)
τ membrane tortuosity
viscous flux (Pa)
density (kg/m3)
ε membrane bulk porosity
εi inner layer bulk porosity
α constant
γ water activity coefficient
γL liquid-vapor surface tension (J/m2)
θ angular coordinate
θef effective contact angle
ρk kerosene density (kg/m3)
ΔHv water latent heat of vaporization (J)
ΔP pressure difference at liquid-vapor interface (Pa)
ΔTp difference between the permeate inlet and outlet temperatures (ºC)
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Chapter 1. Background Review and Objectives
1.1 Introduction
As time goes by, the fresh water shortage becomes a more troublesome issue for human being.
According to the Global Environmental Outlook report by United Nations in 2000, water
shortage, besides the global warming, has been considered as the most worrying problem for the
new millennium [1]. The fresh water shortage is probably related to the rapid increase of human
population and scarcity of the fresh water resources. The following figure schematically shows
the distribution of water on earth.
Fig.1.1 Distribution of Earth’s water [2]
According to the above figure, oceans contain 97% of the world’s water by volume, which is too
salty for drinking. The salinity of oceans is due to the gradual accumulation of dissolved
chemicals eroded from the earth’s crust and washed into the sea by rivers. The average salinity of
sea water is 3.5 wt%, but concentrations as high as 4.0 wt% are observed in the Red Sea and the
Persian Gulf [3]. From the remaining 3% of earth water resources which is considered fresh
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water, 98.8% is trapped in polar ice caps and ground water. Thus, only about 0.03 % of earth’s
total available water by volume is available for human use. Water shortages are expected to
become even more sever in the future due to increasing demand of water and water pollution
issues. However, the water shortage problem is not uniform among different regions and
countries. Depending on the geographical conditions some regions are already facing a water
crisis. The following map shows the distribution of countries with respective shortage of water in
1995 and 2025.
Fig.1.2 Fresh water stress [4]
The above map indicates that a large area in Middle East, North Africa, India and some parts of
Europe are going to suffer from the water shortage problem in the next twenty years. Therefore,
finding alternative water resources seems to be vital to alleviate the water shortages. In this
respect, water desalination has been considered as an important possible fresh water resource for
a long time.
1.2 Desalination Processes
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The main challenge against desalination is to deal with the required energy imposed by basic
thermodynamics for separating salt and water. Theoretically, the minimum energy for
desalinating seawater is around 2.7kJ/kg [5, 6]. In practice, it is not possible to operate processes
at the theoretical minimum energy. Therefore, there is always a trade-off between energy use,
production capacity, and capital costs for optimized industrial processes [7, 8]. Three basic
approaches have been considered for separating water from salt [9]. The first approach is to
induce a phase change in water (to vapor or solid) through heating or cooling, separate the pure
new phase from the remaining salt solution, and then recover the thermal energy for reuse as the
vapor or solid transform to liquid. The second approach is based on the application of semi-
permeable membranes as filters to retain the salt, as saline water is forced through the membranes
due to an externally applied chemical potential gradient. The chemical potential driving force
may be induced by pressure, concentration or electric fields. Finally, there are chemical
approaches to desalination, including ion exchange, liquid-liquid extraction, and gas hydrate or
other precipitation methods [9]. All desalination processes developed or proposed to date have
employed one or more of these three approaches.
Within the three basic approaches to desalination, some of methods have been commercialized on
a large scale; which are mainly distillation and membrane processes. Distillation processes,
including multi-stage flash (MSF), multiple effect evaporation (MEE), and thermal or mechanical
vapor compression (VC), have been used on a large scale for decades, particularly in the Middle
East, and by 1998 accounted for almost 48% of the world’s desalination capacity [10].
Distillation suffers from the inherent drawback of the very large latent heat of vaporization of
water (2200 kJ/kg). The evaporation and condensation steps in these processes are often coupled
so that the latent heat is recovered for reuse by preheating the feed water. Normally, large
distillation units are coupled with steam or gas turbine power plants for better utilization of the
fuel energy. The efficiency and energy consumption of these processes depends largely on the
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effectiveness of recovery of this latent heat in the condensation step. The following figure
schematically demonstrates an MSF desalination plant [10].
Steam heater
Vacuum System
Primary steam source
Flash chambers
Concentrate Out Heat rejection stages
Product Out
Condensate collection trays
Fig.1.3 Schematic presentation of a Multi-Stage Flash desalination plant [11]
In large scale thermal processes, energy consumption has been reported from about 20kJ/kg (VC,
MEE) to more than 200kJ/kg (MSF) of fresh water produced, actually independent of salt
concentration [12]. Such high energy consumption rates have made thermal processes too costly
and energy inefficient.
On the other hand, large scale membrane desalination plants have found great applications during
the last 30-40 years. Three different water resources including seawater, brackish water
(groundwater and surface water), and tertiary treated wastewater can be treated by membranes for
desalination purposes. The most commonly used membrane process for desalination is reverse
osmosis. Reverse osmosis is a separation process that applies pressure higher than a liquid
osmotic pressure to force it through a membrane that retains the solute on one side and permits
the pure liquid to pass to the other side of the membrane. This is the reverse of the normal
osmosis process in which solvent from an area of low solute concentration naturally moves
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through a membrane, to an area of high solute concentration, when no external pressure is
applied. The pressure exerted on the high concentration side of the membrane, is usually in the
range of 2–17 bar (30–250 psi) for fresh and brackish water, and 40–70 bar (600–1000 psi) for
seawater, which has around 24 bar (350 psi) natural osmotic pressure which must be overcome.
Energy requirements for the reverse osmosis process are highly dependent on concentration, and
range from about 10kJ/kg for brackish water and to about 20kJ/kg for seawater; however, new
low energy seawater RO membranes are expected to consume as little as 7.2kJ/kg [10]. The
following figure schematically demonstrates the capacity percentage of desalination plants
worldwide by 1998 [13].
Reverse Osmosis
42%
Electrodialysis 6%
Multi-Effect Distillation
4%
Vapor Compression
4%
Multi-Stage Distillation
44%
Fig.1.4 Schematic diagram of world’s desalination plants capacity percentage by 1998
Compared to energy consumption figures, it is more difficult to estimate cost figures for a
desalination plant. This is mainly due to the fact that many costs, particularly energy costs,
depend greatly on factors such as time, water quality and geography. Also, the transportation cost
of the water to the treatment or distribution point highly depends on location, similar to the cost
of disposing of the resulting concentrate solution [14]. Taking all these limitations into account,
total water costs of $0.75 - $1.5/m3 for thermal processes, vs. $0.25 - $0.70/m3 for brackish water
RO and $0.45 - $1.25 m-3 for seawater RO have been estimated based on a recent review of the
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major current desalination processes [12]. Mainly because of the high energy consumption of
thermal processes compared to RO, membrane processes currently have an enormous economic
advantage.
Fig.1.5 Typical Costs for a Reverse-Osmosis Desalination Plant [15]
1.3 Alternative desalination processes
Despite the success of current technologies, a considerable energy and time is devoted to research
and development of alternative desalination processes. The main driving forces for these efforts
are the need to improve the productivity and the efficiency of a process as well as the desire to
decrease costs. For example, membrane fouling is becoming widely accepted as the largest cause
of permeate flux decline at normal operating pressures and temperatures in brackish water
systems [9, 12]. Furthermore, as the oil price goes up, the operation cost of RO systems increase
very fast. Therefore, alternative desalination approaches offering lower energy consumption or at
least lower energy costs of capital equipment or operation and maintenance are of great interest
among researchers.
Many of the alternate processes proposed in the literature focus on utilization of low-grade
alternative energy sources such as waste heat from other industrial processes like power
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generation plants, solar energy and thermal or mechanical energy from ocean [12]. Among these
approaches, membrane distillation is considered as an emerging desalination alternative providing
several promising advantages over the conventional RO and distillation technologies. Although
energy consumption in this process is quite high, the process is typically run at relatively low
temperature (40 ºC-50 ºC) and thus can make use of waste heat or other relatively low grade heat
sources [12].
1.4 Introduction to Membrane Distillation as an alternative desalination approach
For the first time, membrane distillation was introduced through a US patent by Bodel in 1966
[16]. In this process, a micro-porous hydrophobic membrane is brought into contact with an
aqueous heated solution on one side of the membrane (feed side). The hydrophobic nature of the
membrane prevents penetration of liquid stream through the membrane and creates a vapor liquid
interface at the entrance of each pore. Here, volatile compounds (typically water) evaporate,
diffuse and/or convect across the membrane and condense or remove on the opposite side
(permeate side) of the system [17-20].
Fig.1.6 Schematic diagram representing the separation mechanism involved in MD [19]
1.5 Membrane Distillation Configurations
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Depending on the method used to induce the vapor pressure gradient across the membrane, MD
can be classified into four configurations [17-19]. The most common arrangement is known as
Direct Contact Membrane Distillation (DCMD) in which a condensing fluid stream (often pure
water) is applied in the permeate side of the membrane that is directly in contact with membrane.
Therefore, the temperature difference across the membrane induces a vapor pressure gradient
which is considered as the required mass transfer driving force. In another configuration called
Air-Gap Membrane Distillation (AGMD), a gap of stagnant air is introduced between the
membrane and a condensation surface. Therefore, water vapor molecules have to penetrate
through both the membrane and the air gap to finally condense over a cold surface inside the
membrane module. In Sweeping Gas Membrane Distillation (SGMD) arrangement, the water
vapors diffused to the membrane permeate side are entrained into a condenser by a sweeping gas
such as air and finally pure water is collected through the condenser. In Vacuum Membrane
Distillation (VMD), vacuum is applied in the permeate side of the module by means of a vacuum
pump. The applied vacuum pressure is lower than the saturation pressure of volatile molecules to
be separated from the feed solution. Similar to SGMD, condensation occurs in a separate
condenser outside the membrane module [17-19]. The following pictures schematically illustrate
these arrangements.
Fig.1.7 Schematic diagrams representing different configurations of the MD process [19]
1.6 Temperature Polarization Phenomenon
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As a result of water evaporation and heat conduction and convection from the hot feed side of the
membrane to the cold permeate side, the temperature on the membrane feed side surface is lower
than the feed bulk temperature. The similar phenomenon occurs in the permeate side in case of a
DCMD configuration in which the temperature on the membrane surface will be higher than the
permeate bulk temperature as a result of water vapor condensation; heat conduction and
convection from to the membrane permeate side. Therefore, the mass transfer driving force
(water vapor partial pressure difference at two sides of the membrane) which depends on the
membrane surface temperatures at two sides will be lower. This phenomenon is known as
temperature polarization and leads to flux reduction across the membrane [17-19].
Fig.1.8 Schematic diagram of temperature polarization phenomenon in DCMD
1.7 Wetting Phenomenon
According to the described MD mechanism in the previous section, the hydrophobic membrane
pores are supposed to be dry or in other words occupied by water vapors only. However, if the
liquid stream penetrates into the membrane pores or vapor condensation occurs in the membrane
matrix wetting phenomenon occurs. Depending on the membrane material, pore size and surface
morphology, there is a critical penetration pressure above which the liquid stream on the feed or
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permeate side will penetrate into the membrane [17-19]. This pressure is known as the liquid
entry pressure of water (LEPW) and usually calculated through the following Laplace-Young
equation:
(1.1)
where ΔP is the pressure difference at the liquid-vapor interface, γL is the liquid-vapor surface
tension, θef is the effective contact angle and r is the pore radius.
efp
L
rP
cos
2
1.8 Applications of Membrane Distillation
Based on the mechanism governing the separation process in MD, this process is mostly suitable
for separation of non-volatile components from more volatile liquid solvents such as water [17-
19]. In this respect, desalination has been considered as the mostly used application for MD.
However, MD has also been applied for the environmental waste treatment such as removal of
trace organic compounds from water such as benzene, chloroform and thrichloroethylene [21,
22]. Food processing such as milk and juice concentration is the other application area for MD in
the literature [23, 24]. Other applications of MD include separation of azeotropic mixtures such as
alcohol-water [25, 26], treatment of humic acid solutions [27, 28] and treatment of waste water
contaminated with dyes [29].
1.9 Membrane Distillation Advantages and Drawbacks
The evaporative nature of membrane distillation process makes it possible to achieve an almost
complete rejection of non-volatile solutes, such as macromolecules, colloidal species and ions
through this process. Compared to conventional distillation columns, membrane distillation units
offer a much larger mass transfer area per unit volume which is the characteristic of membrane
based processes. This translates into more compact units with similar fresh water production
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capacities. In addition, the operating temperatures in MD can be maintained as low as 30-50 °C,
thus permitting the efficient recycles of low grade or waste heat streams, as well as the use of
alternative energy resources such as solar power. Furthermore compared to reverse osmosis
process, this process does not suffer limitations of concentration polarization and high operating
pressures. As a result, this process is less susceptible to fouling compared to RO [17-19].
However, despite the distinct advantages MD offers, this process has not been yet implemented in
industry. The main barriers to commercial implementation of this process include [18]:
1- Relatively low permeate flux compared to other separation techniques such as reverse
osmosis
2- Permeate flux decay due to membrane fouling and total or partial pore wetting.
3- Uncertain energy and economic costs for each MD configuration and application
Chapter 2. Literature Review
2.1 Overview on MD literature
The research in MD literature can be divided into different categories based on the area of the
research conducted by researchers during the last 30 years. The initial attempts in 1980s mainly
involved experimental studies using commercial membranes combined with modeling the process
[30-33] for the following goals: 1) to better understand the physics underlying the process 2) to
elaborate the complex heat and mass transfer involved and 3) to estimate the water vapor flux in
the process and investigate the effect of different operating parameters such as permeate and feed
flow rates on the obtained flux. After 30-40 years, it seems that this area of MD research has
become mature and a good understanding of the physics involved in the process has been
achieved. In addition, the theoretical models proposed in the literature can predict the flux with a
high certainty. Another research trend can be identified as the attempt to suppress the undesirable
temperature polarization in the process by fabrication of novel membrane modules, applying
spacers and static mixers in order to induce a better flow geometry inside the module. However,
most of the studies in this area have been related to MD using flat sheet membranes [34-38] and
there are relatively fewer investigations using hollow fibers [39]. Finally, some published reports
can be grouped as the attempts to fabricate novel membranes with higher flux and more
durability. However, the number of published reports on fabrication of MD membranes is much
fewer compared to the previous research trends addressed previously [18]; while many facts
regarding the effects of membrane properties such as morphology, pore size distribution and
thickness on overall MD performance are still unknown. As a result, to the author’s opinion, it is
the time to devote more experimental efforts on fabrication of high performance membranes
optimally designed for the MD process. In this respect, the trends of this research have been
chosen to cover the gaps identified above to some extent.
12
As a result, it is the main objective of this research to enhance the flux in MD process by
fabrication of novel high performance membranes. In this respect, we have devoted this chapter
to review membrane distillation literature regarding membranes applied in this process and their
characteristics.
2.2 Literature review on membrane fabrication for MD
The membranes commonly applied in membrane distillation literature have been the
commercially available membranes actually fabricated for other membrane applications such as
micro-filtration and ultra-filtration [40]. Since the hydrophobic character of the membrane
represents a crucial requirement in MD, membranes have to be made of hydrophobic polymers
having a low surface energy. The most popular hydrophobic polymers applied in the MD include
Polytetrafluoroethelyne (PTFE), Polypropelene (PP) and Polyvinyldene flouride (PVDF).
PTFE having a low surface energy of 9.1 × 10-3 is a highly crystalline polymer which shows
excellent thermal stability and chemical resistance. PP similar to PTFE has a highly crystalline
structure but has higher surface energy compared to PTFE (30.0 × 10-3 N/m). These two polymers
have very low solubility practically in all common solvents. Therefore, micro-porous membranes
are fabricated out of these polymers usually by sintering and stretching methods rather than phase
inversion approaches [41].
In a sintering approach, powders of polymeric particles are pressed into a film or plate and
sintered just below their melting point. This process yields to micro-porous structures having
porosities in the range of 10-40% and a rather irregular pore size distribution. The typical pore
size determined by the particle size of the sintered powder ranges from 0.2 to 20μm [41].
13
In a stretching approach, films are obtained by extrusion of polymeric materials at temperature
close to their melting point coupled with a rapid draw down. Crystallites in the polymer are
aligned in the direction of drawing; after annealing and cooling a mechanical stress is applied
perpendicularly to direction of drawing. This manufacturing process gives a relatively uniform
porous structure with pore size distribution in the range of 0.2-20 μm but it is difficult to obtain
porosities higher than 80% [41].
Fig.2.1 Schematic diagram representing a typical melt spinning process
On the other hand, PVDF is a semi-crystalline polymer with a surface energy of 30.3 × 10-3 N/m
and very good thermal and chemical resistance. This polymer can easily be dissolved in common
solvents such as N-Methyl-2-Pyrrolidone (NMP), Dimethylacetamide (DMAC) and
Dimethylformamide (DMF). Therefore, PVDF membranes can be fabricated through phase
inversion process through which by controlling the fabrication parameters such as polymer
solution concentration, type of coagulants and additives, membranes with higher porosity and
14
narrower pore size distribution can be achieved compared to sintering and stretching approaches
[41].
Although most of the studies in MD (as shown in Table1) have used commercial membranes
originally fabricated for micro-filtration, recently a number of studies have produced their own
membranes in order to improve the flux and separation in this process. There is considerable
number of studies in the literature regarding the fabrication of PVDF membranes applied for
ultra-filtration process [42-45]. However, during the last 10 years, researchers have started to
investigate on the fabrication of PVDF membranes specifically applied for membrane contactor
processes for CO2 capture [46-48], H2S capture [49, 50], VOC removal [51] and desalination
through vacuum and direct contact membrane distillation processes. Oritz de Zarate et al. [52]
fabricated asymmetric flat sheet PVDF membranes through phase inversion approach using
DMAC and DMF as the solvents. They observed that porosity and pore diameter increased as the
PVDF concentration decreased. In another study, Tomaszewska [53] fabricated flat sheet PVDF
membranes using lithium chloride (LiCl) as an additive to the casting solution. The additive LiCl
increased the porosity and pore size of the fabricated membranes. This led to some flux
enhancement in the consequent MD experiments. In a similar study, Khayet and Matsuura [54]
found that the pore size and porosity of the membrane increase with increasing the concentrations
of pure water as a non-solvent into the dope solution. Also, they concluded that the MD flux
increases exponentially with the water content in the PVDF casting solution.
Fabrication of hydrophilic-hydrophobic membranes for MD has been another area of interest for
researchers. For the first time, Cheng and Wiersma [55] described the use of composite
membranes in MD in a series of patents. They modified a cellulose acetate membrane via
radiation graft polymerization of styrene onto the membrane surface, and a cellulose nitrate
membrane via plasma polymerization of vinyltrimethylsilicon/carbon tetrafluoride and
15
octafluoro-cyclobutane. Wu et al. [56] applied hydrophilic porous supports such as cellulose
acetate, and treated the membrane surface via radiation graft polymerization of styrene to enhance
the hydrophobicity. In a similar way, Kong et al. [57] modified a cellulose nitrate membrane via
plasma polymerization of both vinyltrimethylsilicon/carbon tetrafluoride and octafluoro-
cyclobutane.
Such composite membranes have been considered to be applied in MD in two different
configurations. In the first arrangement, a thin hydrophobic functional layer is supported by a
rather thick hydrophilic layer so that the water vapor mass transfer resistance through the
hydrophobic layer is minimized and a greater flux can be obtained. Fig.2.2 demonstrates this
concept schematically.
In the literature, Khayet and Matsuura, and their coworker [58, 59] fabricated flat sheet
hydrophilic - hydrophobic membranes for this purpose. They fabricated such composite
membranes based on the migration of hydrophobic macromolecules (SMM) to the membrane
surface. However, the flux obtained through their experiments was very low. This might be due
to the difficulty to optimize the membrane surface porosity and pore-size distribution in this
process.
Hydrophobic layer
Hydrophilic layer
Fig. 2.2 A composite hydrophilic-hydrophobic membrane before the MD test (left) A composite hydrophilic-hydrophobic membrane during the MD test (right)
16
The second area that the application of composite hydrophilic-hydrophobic membranes may be
promising is the fouling and wetting prevention of the hydrophobic functional layer by coating a
very thin hydrophilic layer in contact with the feed salt solution. The hydrophilic layer is thought
to be less susceptible to fouling and scaling phenomena. In this respect, Peng et al. [60] tested the
desalination through DCMD process by applying a composite PVA/PEG hydrophilic layer on a
hydrophobic PVDF substrate. The authors investigated the effects of brine temperature, salt
concentration, running time and the addition of ethanol on the flux of composite membranes.
They observed that durability of the membranes greatly improved; while the flux obtained
through the process did not change greatly compared to the flux in case of the single layer
hydrophobic membrane.
Among the fabrication methods applied in the literature, both radiation graft and plasma
polymerizations are expensive processes that limit their applications. On the other hand, as
pointed out above, fabricating these types of membranes based on the hydrophobic
macromolecules migration to the surface suffers from the difficulty to optimize the hydrophobic
layer thickness and morphology. In addition, all the related reports in the literature so far have
been limited to the flat sheet membranes while hollow fiber is the most preferable membrane
configuration for the industry because of providing high surface area per unit volume as well as
ease of module fabrication.
2.3 Research objectives
In overall, despite the attempts that have been initiated to fabricate high performance membranes
for the MD process, the obtained water vapor fluxes through the reported studies scatter greatly
(0.2-40 kg/m2hr) and are quite low in average. As a result, the objective of this research is to
develop novel high performance membranes and optimize their different separation properties in
17
order to enhance the MD flux and at the same time maintain a high salt rejection rate and stable
long term performance of the membranes. It is also worthy to note that the focus of this research
will be on hollow fiber membranes mainly because they have been considered as the most
suitable membrane configuration for the industrial applications. Hollow fibers provide a high
surface area per unit volume and ease of module fabrication. In the following chapter, we
demonstrate the fabrication of dual layer hydrophilic-hydrophobic hollow fibers through a novel
co-extrusion approach.
Fig.2.3 Schematic picture of a dual layer hydrophilic-hydrophobic fiber
18
Chapter 3. Theory and model development
In this chapter we review the concepts of heat and mass transport phenomena involved in the MD
process and later we develop a mathematical model for the DCMD process in a hydrophobic-
hydrophilic hollow fiber membrane to fit our experimental flux data in the next chapters. Finally,
we discuss the characteristics of a high performance dual layer hydrophilic-hydrophobic fiber
based on the theoretical concepts.
Fig.3.1 Schematic DCMD process with dual layer hydrophilic-hydrophobic hollow fibers
3.1 Mass transfer
Water vapor transport through hydrophobic membranes in MD process can be considered as the
mass transfer within the gas-phase of a porous medium. Depending on the pore size of the porous
medium, continuum and Knudsen regions can be identified in the membrane matrix [19, 61]. In
the continuum region, the mean free path of the gas is small compared with the pore diameter. In
19
this region, molecule-molecule collisions are dominant over molecule-wall collisions and the
mass transfer flux can be described by the Fick’s law. On the other hand, in a Knudsen region the
mean free path of the gas is larger than the pore diameter. As a result, molecule-wall collisions
are dominant over molecule-molecule collisions and the mass transport can be described by
Knudsen’s law. Transition region between continuum and Knudsen regions may also be
identified when the mean free path of gas molecules is in the same order as the membrane pore
diameter. Dominance, coexistence, or transition between all of these different mechanisms can be
quantitatively estimated by dimensionless Knudsen number which is defined as the ratio of mean
free path of diffusing molecules to the mean pore size of the membrane. According to the kinetic
theory of gases, the mean free path of molecules can be calculated through [19, 61]:
22
P
TkB (3.1)
where kB is the Boltzmann constant (1.380 × 10-23 JK-1), and σ is the collision diameter of the
molecule (σ =2.7Aº for water). For the binary mixture of water vapor in air, the average free
mean path of air-water molecules mixture can be evaluated at an average operating temperature
through the following equation [19]:
a
wwa
Bwa
M
MP
Tk
12
2,
(3.2)
where σa (=3.7 A° ) and σw are the collision diameters, and Ma and Mw the molecular weight for
air and water, respectively. For an average temperature T=60 ºC, Phattaranawik et al. [62]
reported a mean free path of 0.11µm that is in the same order as the pore size of the micro-porous
membranes commonly applied in the MD process. As a result, the transition region between
Knudsen and ordinary molecular diffusion is normally considered as the governing mass transfer
mechanism for the MD membranes.
20
Accordingly, the flux relationship in a DCMD process has been given in the literature as follows
[19, 61].
wa
ww
w
aw
a
wk
ww PP
PN
dr
dP
D
P
DRT
MN
11/
(3.3)
where Nw is the water vapor mass flux, Mw is the water molecular weight, R is the universal gas
constant, Pw and Pa are the water vapor and air partial pressures, respectively, r is the radial
coordinate, ε and τ are the membrane porosity and tortuosity, respectively, T is the temperature,
Dwk and Dw-a are the water vapor Knudsen and molecular diffusion coefficients, respectively.
According to this equation, the flux depends on the partial pressure of air trapped inside the
membrane, which brings some uncertainty into the modeling practice as the air partial pressure
profile depends on the initial amount of air trapped in the membrane matrix, which itself varies
upon the operating conditions such as feed and permeate hydrodynamic pressure and also the start
up procedure as well. In addition, there is no practical approach to determine the tortuosity factor
exactly. In order to avoid these complexities, researchers have often applied the following
simplified flux equation with a lumped parameter Ce which is fitted into the experimental data.
[19]
dr
dP
RT
MCN ww
ew (3.4)
Using a similar approach and assuming a steady state and one dimensional mass transfer, the
following equations and boundary conditions can be applied:
01
wrNdr
d
r ,
dr
dP
RT
MCN ww
ew (3.5)
ww xTC
BAP
exp at oRr (3.6)
TC
BAPw exp at mRr (3.7)
21
where A, B, C are Antoine constants, γ is the water activity coefficient in an aqueous sodium
chloride solution, Ro and Rm are the outer and inner radii of the hydrophobic layer, respectively
and xw is the mole fraction of water in the feed.
3.2 Heat transfer
Beside mass, heat is also transferred from the feed to the permeate side by conduction through the
membrane matrix as well as convection of water vapors across the membrane. As a result of water
vaporization and heat conduction at the membrane feed side, the membrane surface temperature at
this side will be lower than the feed bulk temperature. On the other hand, because of water vapor
condensation and heat conduction at the permeate side, the membrane surface temperature at this
side will be higher than the permeate bulk temperature. This phenomenon is known as temperature
polarization which lowers the temperature and vapor pressure driving forces across the membrane
and consequently leads to a lower flux. The following equations describe the explained heat
transfer for a steady state and one dimensional process mathematically [63].
dr
dTrk
dr
d
rdr
dTCN vw
1 (3.8)
ffvw TThd
r
dTkHN at oRr (3.9)
ppvw TThd
r
dTkHN at mRr (3.10)
where k is the membrane hydrophobic layer thermal conductivity, Cv is the water vapor heat
capacity at the constant volume, hf and hp are the feed-side and permeate-side heat transfer
coefficients, respectively and ΔHv is the water enthalpy of vaporization. As heat is transferred
between two sides of the membrane, the feed temperature decreases while permeate temperature
increases along the membrane module. This can be expressed in terms of equations as follows
[63]:
22
TThnATCmdz
dfffofff at r at oRoRr (3.11)
fif TT at 0z (3.12)
TTHnATCmdz
dpppoppp at mRr (3.13)
pip TT at 0z (3.14)
(3.15)
i
mi
im
p
p
R
Rk
RR
h
H
ln
11
ph
1= convection resistance at permeate side
i
mi
im
R
Rk
RR
ln
=hydrophilic layer resistance (3.16)
where Hp is the overall heat transfer coefficient in the permeate side, ki is the thermal conductivity
in the hydrophilic layer,
and are the feed and permeate mass flow rates, respectively,
while Cf and Cp are the feed and permeate heat capacities, respectively. n is the number of hollow
fibers in the membrane module, Afo and Afi are the hollow fiber outer and inner surface areas,
respectively. Tfi and Tpi are the feed and permeate inlet temperatures, respectively.
fm pm
In order to estimate the flux obtained through the process, the following procedure was adopted:
First, the feed and permeate heat-transfer coefficients were estimated using the existing
correlations in the literature. For the laminar permeate flow inside the fibers the following
equation was applied [64].
Nu=1.86 (di/L Red Pr)0.33 (3.17)
Nu=hp di/kw (3.18)
23
where di is the fiber inner diameter, L is the fiber length, Red is the Reynolds number based on
fiber inner diameter as the characteristic length and Pr is the Prandtl number. The heat-transfer
coefficient in the module shell side was estimated using the following equation, which is
correlated by Mengual et al. [65] for an external parallel flow along the fibers.
Nu=0.042 Re0.59 Pr0.33 (3.19)
Nu=hf L/kw (3.20)
The membrane thermal conductivity was estimated using the following equation which has
commonly been applied by researchers in membrane distillation literature [17, 19].
k = kair ε + km (1-ε) (3.21)
where k is the effective membrane thermal conductivity, kair and km are the air and polymer matrix
thermal conductivities, respectively The describing equations (3.3 - 3.21) were solved using an
iterative finite difference scheme and the membrane diffusivity coefficient Ce was obtained by
fitting the calculated flux to the experimental data at different feed inlet temperatures.
3.3 Characteristics of a high performance MD membrane
3.3.1 High membrane permeability
Based on the mass transfer mechanism in MD membranes that was described in the previous
section, highly gas permeable membranes are essential to achieve a high flux through the process.
In this respect, membrane porosity has been known to be a key parameter influencing the
membrane gas permeability [17-19]. According to the equations (4) and (5), as the membrane
bulk porosity increases the molecular and Knudson diffusivities of the vaporized molecules
diffusing through the membrane increase in a proportional manner. Furthermore, as the
membrane bulk porosity increases, the membrane thermally conductivity decreases and leads to a
24
lower heat loss and a higher energy efficiency of the membranes. The energy efficiency is usually
defined as the fraction of feed brine thermal energy that has been used to evaporate water. Energy
efficiency can be easily calculated using the following equation [64]:
(3.22)
Where J is the water vapor flux, ΔHv is water vapor latent heat of condensation, Ao is the total
membrane area based on the fibers’ outer diameter, is the permeate mass flow rate, Cp is the
specific heat capacity of the permeate stream and ΔTp is the difference between permeate inlet
and outlet temperatures. This is due to the fact that a higher volume percentage in the membrane
will be occupied by the less thermal conductive air molecules (kair=0.024 W/m°K) compared to
the polymer matrix (km=0.1-0.3 W/m°K).
pm
ppp
ov AHJEfficiencyEnergy
TCm
It is also worthy to note that besides membrane bulk porosity, membrane surfaces must be highly
porous as well, so that a high liquid-vapor interfacial area is formed at the membrane pore
mouths. In other words, a membrane with a higher surface porosity provides a higher effective
mass transfer area which translates into a greater trans-membrane flux [66].
In addition to the membrane porosity and pore size, membrane morphology and pores
interconnectivity are other determining factors that have been never discussed in the MD
literature. However, there are a number of studies regarding the influence of these parameters in
other membrane separation processes. Li and Chung [67] identified pore interconnectivity to have
a great influence on the pure water permeability (PWP) of SPES hollow fibers applied in ultra-
filtration. In other words, between two membranes with similar porosities, the membrane with an
open cell structure or a high degree of pores interconnectivity will possess a higher permeability.
In another study, Widjojo et al. [68] found that high pores interconnectivity in the supporting
25
layer of the 6FDA-ODA-NDA/PSF dual-layer hollow fiber membranes could greatly reduce the
substructure resistance in gas separation applications. Therefore, high porosity, big pore size and
high degree of pores interconnectivity should be maintained in order to maximize the membrane
permeability.
3.3.2 High membrane wetting resistance and long term stability
As discussed in chapter 1, membrane wetting is an undesirable phenomenon in MD that may lead
to flux decay or lower separation efficiency. In order to prevent this phenomenon in long term
MD operations, high membranes liquid entry pressure, proper pore size distribution (0.2-0.5 µm)
[69] and a reasonable degree of membrane tortuosity are essential.
In term of membrane permeability, macrovoids may seem to form desirable membrane structures
with low tortuosity and high porosity. But from a long term performance point of view,
macrovoids might be undesirable structures that enhance the wetting phenomenon in membranes,
consequently reduce the membranes long term stability. Therefore, fabrication of macrovoid-free
and fully sponge-like membranes with sharp pore size distributions is considered as the other
characteristics of an ideal MD membrane.
3.3.3 Suitable membrane geometry and dimensions
Similar to other membrane processes, membrane thickness has a significant effect on the overall
mass transfer resistance in the MD process [18]. However, in contrast to other membrane
approaches the membrane flux does not have a monotonically increasing trend as function of
membrane thickness [70, 71]. This is due to the fact that flux in MD depends on the coupled
effect of mass and heat transport phenomena. On one hand, a smaller membrane thickness
decreases the mass transfer resistance introduced by the membrane and leads to a higher flux. On
the other hand, as the membrane thickness decreases, heat loss through the membrane increases
26
which translates into a lower temperature gradient across the membrane, a lower driving force
and ultimately a lower flux [70, 71]. Therefore, there should be an optimum thickness for the
membranes depending on their other properties as well as the flow geometry. One way to tailor
the membrane thickness to the optimum range is to fabricate hydrophilic-hydrophobic
membranes in which a thin functional hydrophobic layer is supported by a hydrophilic layer. In
this way, as discussed above, the water vapor mass transfer resistance by the functional
hydrophobic layer thickness will be minimized while the hydrophilic layer will act as a support
for the thin functional layer so that the whole membrane is mechanically strong.
For the fabrication of hydrophilic-hydrophobic membranes with hollow fiber geometries some
other considerations come into practice, which are discussed in the following sections:
3.3.4 Hydrophilic layer porosity and thermal conductivity
In hydrophilic-hydrophobic membranes, the hydrophilic layer introduces an additional heat
transfer resistance that amplifies the undesirable temperature polarization. Based on equation
(3.13), in order to minimize the heat transfer resistance induced by a fixed hydrophilic layer
thickness, one possible way is to maximize the thermal conductivity of the hydrophilic layer ki by
fabricating the layer as porous a possible. This can be understood according to the following
equation describing the inner layer thermal conductivity.
27
imiwi kkk 1 (3.24)
Therefore, as the porosity of the hydrophilic layer increases, a higher volume percentage of the
hydrophilic layer will be occupied by the more conductive liquid water (kw = 0.58 W/m°K > km =
0.1-0.3 W/m°K) during the DCMD operation. Hence, the thermal conductivity of the layer
increases accordingly.
Chapter 4. Experimental
4.1 Materials
PVDF polymer was chosen to be applied as the main polymeric material to form the functional
membrane applied in the MD process. PVDF is a highly thermal and chemical resistant polymer
with a high hydrophobicity that corresponds very well to the properties required for a high
performance MD membrane. The highly hydrophobic nature of this polymer can be attributed to
the presence of two Fluor atoms in its monomer chemical structure that lowers the surface tension
of the polymer to about 33.7mN/m (25dynes/cm) [72]. PVDF Kynar HSV 900 was used in our
experiments and was purchased from Arkema Inc.
Fig.4.1 Chemical structure of PVDF polymer
Polyacrylonitrile (PAN) was applied in the inner layer of dual layer fibers in order to form a
hydrophilic support layer for the PVDF functional outer layer. PAN is a highly hydrophilic
rubbery polymer which is usually used in a fiber form as the chemical precursor of high quality
carbon fiber. Homopolymers of PAN have applications as fibers in reinforced concrete, hot gas
filtration systems and outdoor awnings. But mostly PAN copolymers are used as fibers to make
knitted clothing as well as outdoor products such as tents [73]. PAN was kindly supplied by
Profs. J.Y. Lai and H.A. Tsai at Chung Yuan Christian University of Taiwan. N-Methyl
Pyrrolidone (NMP) and methanol were used as solvent and nonsolvent respectively and supplied
by Merck.
28
Fig.4.2 Chemical structure of PAN polymer
Hydrophobic cloisite 15A and hydrophilic cloisite NA+ montmorillonite clay particles were
purchased from Southern Clay Products. These particles were mainly applied as additive to
polymeric dope solutions in order to enhance their physical properties and processability.
Furthermore, they were targeted to amend surface tension properties of the inner and outer layers
of membranes and also to create more free volume and pores in the membrane structures. Table.1
shows some properties of these particles provided by the supplier.
Typical dry particle size Treatment-
properties
Organic
Modifier
Modifier
Concentration 10% less
than
50% less
than
90% less
than
Density Hydrophilicity
Hydrophobicity
Cloisite 15A 2M2HT 125 meq/
100 g clay 2 μm 6 μm 13 μm 1.66 g/cc Hydrophobic
Cloisite NA+ None None 2 μm 6 μm 13 μm 2.86 g/cc Hydrophilic
Table.4.1 Specificcations of Cloisite particles
4.2 Dope preparation
The polymers were dried at 60 °C under vacuum overnight before they were used for dope
preparation. Different concentrations of PAN/NMP, PVDF/NMP and mixtures of these solutions
with clay particles were prepared by stirring the solutions at 65 °C for solutions containing PAN
and ambient temperature for other solutions, over a period of 12 hours. The prepared solutions
were degassed overnight before they were used for spinning.
29
4.3 Fabrication of flat sheet membranes
In order to investigate the effect of different coagulants on the surface porosity and roughness of
the PVDF membranes, PVDF flat sheet membranes were cast with a thickness of 250 μm using a
knife casting approach on a glass plate. Afterwards, the films were soaked into coagulation bath
containing pure water, 40/60 and 80/20 wt. % methanol and water for 24 hours. Consequently,
the cast films were freeze-dried. The fabricated films were used for SEM surface morphology and
contact angle measurement tests
4.4 Fiber spinning
Single and dual layer hollow fibers were spun through the spinning line schematically shown in
the following figure.
Fig. 4.3 Schematic representation of the fiber spinning line
At least 8 hours before starting the spinning, polymer solutions were doped into syringe pumps so
that a complete degassing of dope solutions is achieved. On the other hand, another syringe pump
was filled with bore fluid and the syringe pumps were connected into specific channels of the
spinning device called spinneret. Spinning initiated by turning on the pumps and choosing
desirable dope and bore fluid flow rates. The discharged flow form the spinneret went through an
30
air-gap distance ranging from 0-20 cm depending on the experimental conditions and entered an
external coagulant bath where the main phase inversion process occurred. The nascent fibers were
collected using a drum which induced a drawing force and consequently take-up speed on the
fibers. The take up rate ranged from free fall to higher take up rates depending on the
experimental conditions. The as-spun fibers were immersed in water for approximately two days
for solvent exchange. Two different drying procedures were initially applied for each spinning
batch. In the first procedure, fibers were immersed in methanol once for 30 minutes. Afterwards,
the fibers were dried by air at room temperature. In the second drying procedure, fibers were
frozen in a freezer for about 5 hours and then freeze dried using a Thermo Electron Corporation,
Modulyod freeze-dryer for about 8 hours. By comparing the two procedures, we found out that
freeze drying prevents the sever polymer shrinkage, which occurs during air drying. It was
observed that the fiber shortening as much as 50% occurred in case of the fibers dried in air.
Therefore, the freeze drying was adopted as the standard drying procedure for all the experiments.
4.5 Morphology study of hollow fibers by SEM
Hollow fiber cross-sections were observed by taking SEM pictures using a JEOL JSM-5600LV
scanning electron microscope. Fiber samples were immersed in liquid nitrogen, fractured and
then coated with platinum using a JEOL JFC-1300 coater.
4.6 Contact angle measurements
A Sigma 701 Tensiometer from KSV Instruments Limited was used to measure the contact angle
of a hollow fiber as follows: The hollow fiber was immersed into distilled water and the
advancing contact angle was calculated with the aid of computer software. Ten readings were
measured and an average was obtained from the results. In case of hydrophilic-hydrophobic
hollow fibers, in order to prevent the effects of the inner hydrophilic layer on the contact angle
measurements for the outer layer, the lumen side of the fibers was sealed by inserting a metallic
31
bar into the fiber lumen side. Furthermore, in order to estimate the contact angle of the inner
hydrophilic layer, flat sheet membranes were cast out of the same dope solution used for the inner
layer and immersed in a coagulant composition similar to the fiber bore fluid. The same
procedure was performed for the outer layer dope solution in order to confirm whether the contact
angle measurements of flat sheets can give a reliable estimate of the measurements performed for
hollow fibers. The fabricated flat sheet membranes were tested for contact angle measurement
using a Rame-Hart contact angle goniometer (model 100-22) by the sessile drop method at 25 °C.
A built-in image system provided by Rame-Hart was able to acquire the image, transmit to a
computer, and perform the image analysis.
4.7 Porosity measurement
Two different approaches were adopted for hollow fibers porosity measurements. In the first
approach, porosity measurements were carried out using an Autopore III 9420 mercury
porosimeter. A hollow fiber sample was cut into about 20 pieces, each piece of about 2 cm length.
Then, the fiber pieces were loaded into a penetrometer and the penetrometer was loaded and
installed in a low pressure port. At the first phase of the low pressure analysis, the penetrometer
was evacuated from gases. Later it was automatically backfilled with mercury and data were
collected at pressures up to about 50 psia. When the low pressure analysis was complete the
penetrometer was removed from the low pressure port and installed in a high pressure port. At the
high pressure analysis, data were collected at pressures up to 60,000 psia.
In the second approach, the sample hollow fibers were first weighed with a beam balance (w1),
following by immersion in 33% LIX54 kerosene solution for one week. An assumption was made
where all the empty voids were filled with the liquid membrane phase. Next, the fully
impregnated fibers were removed from the kerosene and any excess kerosene in the lumen side
32
and on the outer surface of the fibers was wiped away. Afterwards, the fibers were weighed
again. The empty volume (V) can be calculated by the following relation:
k
wwV
12 (4.1)
where ρk is the kerosene density. The porosity of the hollow fibers membrane (ε) was estimated
by the ratio of empty voids (V) to the total volume of the membrane samples [74]. An average
value of the two tests was considered for the fibers porosity.
4.8 Pore-size distribution measurement
Pore size distribution measurements were carried out using a PMI capillary flow poromter. The
sample modules were fabricated in a similar way as we did for gas permeation rest. The
fabricated modules were immersed in a low surface tension liquid such as isopropyl alcohol (IPA)
for one day so that the pores are fully saturated with the liquid. Afterwards, the Wet Up/Dry
Down test was performed using the porometer. In the Wet Up portion of the test, the sample was
wet (saturated with a wetting fluid) and the porometer acquired data points with increasing N2 gas
pressure. On the Dry Down portion of the test, the sample was dry and data points were acquired
as the pressure was gradually decreased.
4.9 Gas permeation test
Pure gas N2 permeation measurements were carried out to test the permeance of the fibers
according to the procedure described in [75]. Gas permeation modules were composed of 2
hollow fibers with a length of around 7 cm. The permeate flow rate was measured at different
pressures using a bubble flow meter. The permeance P/L was measured by the following equation
[76]:
(4.2)
PnDl
Q
PA
Q
L
P
33
where P is the permeability of separating layer (Barrer), L the thickness of the fibers (cm), Q the
pure gas flux (cm3/s), n the number of fibers in one testing module, D the outer diameter of the
testing fibers (cm), l the effective length of the modules (cm), and ΔP the gas pressure difference
cross the membrane (cmHg).
4.10 Polymer flow observation by a high magnification camera
By making use of the transparent nature of PAN dope solution we were able to take high
magnification pictures of the polymer dope and bore fluid flow geometries in the air-gap region.
An EOS 350 digital Canon camera with an MP-E65mm high magnification lens was utilized for
the digital visualization.
4.11 Video microscopy flow visualization
A drop of dope solution was sandwiched between two microscope slides. Afterwards, a drop of
water as a coagulant was introduced to the gap between the two slides by a syringe. The
diffusion, precipitation and solidification fronts of water into the dope solution were observed and
video-recorded under an Olympus BX50 polarizing optical microscope. These experiments may
provide the basic morphological evolution of the phase inversion process.
4.12 Module Fabrication
In order to fabricate membrane modules, hollow fibers were cut in about 25 cm long and placed
parallel to each other and were put together and bundled using a Teflon strip in two ends. The
prepared bundle was inserted inside the fitting and tube assembly as shown in Fig. 3.2.
34
Fig.4.4 Schematic diagram showing the steps involved in fabrication of lab scale membrane modules [77]
In the next step, the space between the fiber bundle and the fittings at the two ends of the module
were filled with hydrophilic cotton. The cotton is applied as a substrate to absorb the epoxy and
form a sealing agent among the fibers in order to separate the shell and module side of the fibers.
Later, the module was placed vertically on a holder and a proper mixture of epoxy resin and
hardener (1 to 2 in this study) was added to the higher section of the module standing on the
holder. After a few hours that the epoxy is completely hardened the same epoxy potting
procedure was carried out for the other side of the module.
4.13 DCMD experiments
DCMD experiments were carried out using the experimental set-up shown in the following
figure.
35
Fig.4.5 Schematic experimental set-up applied for direct contact membrane distillation process
There were two separate feed and permeate cycles in this set-up, which have been highlighted by
red and blue colors, respectively. A centrifugal pump circulated the preheated saline feed water
with 3.5 wt% sodium chloride (NaCl) concentration from the feed tank through the shell side of a
vertical membrane module. A saline solution of 3.5 wt% was used to represent the ocean water
which has an average salinity of 3.5 wt%. On the other hand, a peristaltic pump circulated
permeate pure water from a conical flask on a balance through a cooler and consequently through
the lumen side of the membrane module. The feed and distillate streams flowed co-currently from
the bottom to the upper part of the MD module. The ionic conductivity of the permeate stream
was measured before and after the test using a Radiometer analytical conductivity meter (model
Pionneer30) in order to calculate the separation factor.
36
Chapter 5. Fabrication of dual layer hydrophobic-hydrophilic hollow fibers
5.1 Effect of coagulant on the surface morphology of PVDF membranes
Fig.5.1 shows the SEM pictures from the top surface (facing the coagulant) of the PVDF flat
sheet membranes for different coagulant compositions.
Water/Methanol (20/80 wt.%) Water/Methanol (60/40 wt.%) Water
Fig.5.1 SEM pictures from the top surface (facing the coagulant) of the PVDF flat sheet membranes for different coagulant compositions
It can be clearly observed from Fig. 5.1-A that when a strong coagulant such as water was used
for the fabrication of the membranes, dense and smooth surfaces with no obvious pores was
formed. However, when we introduced methanol (a weaker coagulant compared to water) to the
coagulant up to 40 wt.%, the roughness of the membrane surface increased considerably while the
number and the size of the surface pores increased as well (Fig. 5.1-B). Finally, as it is shown in
Fig. 5.1-C, a coagulant containing 80 wt% methanol induced a porous 3-dimensional fiber-like
surface morphology with some degrees of spherulitical structure. The pore sizes are in the range
of 0.4 μm on the membrane outer surface. The formation of a 3-dimensional fiber-like structure is
probably due to the delayed demixng, spinnodal decomposition and then coarsening [78].
37
Spinnodal decomposition is a phase separation process which occurs in unstable systems. In other
words, the small perturbations in the system are amplified and any small change in system
composition lowers the free energy of the system. This process differs from nucleation and
growth phase separation in a sense that in spinnodal decomposition phase separation
occurs throughout the material rather than nucleation sites only [78].
The difference in morphology may be explainable from the solubility parameters of PVDF, water
and methanol as summarized in Table 3.
Solvent Water
[79]
Water/methanol
(60/40 wt.%)
Water/methanol
(20/80 wt.%)
Methanol
[79]
NMP
[80]
PVDF
[81]
δt (MPa 0.5) 47.8 41.6 35.5 32.4 11.2 23.2
δt (Cal 0.5 /cm1.5) 23.4 20.3 17.4 15.9 5.5 11.3
Table 5.1 Total solubility parameter (δt) of water, methanol, NMP and PVDF
Water is the strongest non-solvent toward PVDF compared to 60/40 wt.% and 80/20 wt.%
water/methanol mixtures. Therefore, the liquid-liquid miscibility gap for PVDF-NMP-water will
be the biggest, while this gap for PVDF-NMP-water/methanol (20/80 wt.%) will be the smallest.
As a result, for the PVDF-NMP-water system, the solution may go through rapid phase separation
via nucleation growth and thus produce a rough dense surface. However, in case of PVDF-NMP-
water/methanol (20/80 wt.%) system, the liquid-liquid miscibility gap is small that favors a delay
demixing (i.e., a slow phase separation). With an increase in methanol and water content diffused
from the external coagulant into the nascent fiber, the precipitate path may go through spinodal
decomposition and form a 3-directional fiber-like network [78]. The apparent spherulitical
structure upon the fiber network may be due to coarsening effects. Therefore, in order to induce
high porosity and roughness on the inner and outer surfaces of the PVDF hollow fibers, a mixture
38
of methanol/water with 80 wt% methanol concentration was adopted as the standard coagulant to
be used as the bore fluid and external coagulant in the spinning line.
5.2 First batch of fiber spinning
In our first batch of spinning, we adopted PVDF and PAN polymers as the hydrophobic outer
layer and the hydrophilic inner layer of the spun fibers respectively. It is known that a dope
solution with a low polymer concentration undergoing a non-solvent induced phase separation
tends to result in more porous membranes [82, 83]. Therefore, low concentration polymer
solutions of PVDF/NMP and PAN/NMP were used for outer and inner layer dope solutions
respectively. The following Table shows the composition of the inner and outer layer dope
solutions used for the spinning of this batch with other spinning conditions.
Table.5.2 Spinning conditions applied for the first batch of spinning
39
5.2.1 Membrane morphology
The following SEM pictures show the cross section and surface morphology of the spun dual
layer fibers.
Fig.5.2 SEM pictures representing the structure of first spinning batch fibers
As can be observed in Fig.5.2, the fabricated fibers showed desirable membrane morphology and
structures without any delamination. The outer hydrophobic PVDF layer was 30 µm thick with a
fully sponge-like structure and high surface porosity. The inner support layer was 150 µm thick
with an asymmetric structure with some macrovoids observable in the inner contour of the fibers.
40
5.2.2 DCMD performance
Although the fibers in the first spinning batch were fabricated in a desirable structure, when they
were applied in the DCMD separation process, the interface between the two layers was pealed
off and no separation could be achieved. This might be due to different thermal expansion
coefficients of these two polymers or the repulsive force acted on the interface between
hydrophilic and hydrophobic layers in an aqueous environment.
Fig.5.3 Delamination phenomenon during the DCMD process using the first batch of fibers
In addition, in other attempts to fabricate single layer PVDF hollow fibers using weak coagulants
such as a methanol-water mixture containing 80 wt.% methanol, we noticed that the nascent
fibers formed in the spinning line were mechanically weak and the spinning process was
accompanied with instabilities, fiber breakage and fiber collapse in the collecting wheel.
5.3 Second batch of hollow fiber spinning
In order to overcome the obstacles we faced during the first batch o spinning, we formulated new
dope solutions containing PVDF as the major component in both layers. In this way, we expected
to overcome the problem regarding the interface separation of the two layers containing polymers
with different surface tension and thermal expansion characteristics. However, in order to induce
different surface tension properties in the two layers and simultaneously increase both the
41
processability and mechanical strength of nascent fibers, hydrophilic and hydrophobic clay
particles were added to the inner and outer layer dope solutions, respectively. Clay particles have
been extensively used for polymer reinforcement [84]. Furthermore, some amounts of highly
hydrophilic PAN polymer was added to the inner dope solution in order to enhance the water
transfer from the fiber permeate side into the interface between the two layers. We postulate that
the PAN polymer chains could act as bridges among the hydrophilic clay particles which can be
considered as islands in the polymer matrix.
As a result, modified dope solutions made of 12.5/87.5 (in wt. %) PVDF/NMP containing 30
wt% hydrophobic cloisite 15A loading in the total solid was chosen for the outer layer, and a
dope solution made of 8.5/4/87.5 (in wt%) PVDF/PAN/NMP containing 50 wt% hydrophilic
cloisite NA+ loading in the total solid was selected for the inner layer. The maximum amount of
30 wt% and 50 wt% loadings in total solids were adopted because in higher loadings the
fabricated membranes became rather brittle and the spinneret blocked easily. Less loading was
adopted for the outer layer because of the smaller spinneret discharge gap for the outer layer
compared to the inner layer. The following schematic figure summarizes the spinning conditions
used in the second batch of spinning.
42
Fig.5.4 The spinning conditions applied in the second batch of spinning process
5.3.1. Hollow fibers morphology
The SEM and EDX micrographs in Fig.5.5 show the cross section morphology and clay particle
distribution in the fabricated fibers. It can be observed that a desirable sponge like structure with
rather uniform distribution of particles has been achieved. , the thickness of the outer hydrophobic
layer is about 50μm and the hydrophilic support layer is about 330 µm. On the other hand, as can
be observed in Fig.5.6 a high outer surface porosity has been formed which is in agreement with
our earlier observations for the flat sheet membranes as shown in Fig.5.1, fabricated using a weak
coagulant.
43
Fig.5.5 Cross section morphology of fibers fabricated through the second batch of spinning
Fig.5.6 SEM micrographs showing the surface morphology of the dual layer fibers
44
5.3.2. Pore-size distribution
The pore size distribution of the newly developed hollow fibers was also obtained. As can be
observed in Fig.5.7, a very sharp pore size distribution with an average pore size of 0.41μm has
been achieved. Such relatively big pore sizes with a sharp distribution may lead to water vapor
mass transfer enhancement in the MD process.
Pore Size (µm)
Fig.5.7 Pore size distribution of the fabricated fibers
5.3.3. Gas permeationand porosity measurement tests
A high permeance of 1.25 cm3 (STP)/(cm2)s.cmHg was achieved for the fibers tested for the N2
gas permeation test. Such a high permeance is in agreement with a measured porosity of about
80% and big pore size of the fibers.
5.3.4. Contact angle measurements
The graph in Fig.5.8-A represents the measured force per unit length of the hollow fiber versus
the immersion depth obtained from the hollow fiber contact angle measurements corresponding to
an advancing contact angle of 136°, while Fig.5.8-B shows the equilibrium state of a water
45
droplet on the fabricated flat sheet membrane cast out of the outer layer dope solution using the
same coagulant, which corresponds to a contact angle of 140°. The high contact angles obtained
can be attributed to the high roughness induced on the membrane surface caused by the porous 3-
dimensional fiber-like network with some degrees of spherulitical structure and clay particles
used in the membrane matrix. Since the values obtained from the two methods were quite
comparable, we estimated that the contact angle of the fiber inner layer should be quite similar to
the equilibrium contact angle of the flat sheet film cast out of the inner layer dope solution. A
very low contact angle of 50° was obtained from this experiment, which exhibits the hydrophilic
character of the inner layer (Fig.5.8-C).
Fig.5.8 Contact angle measurements of the fabricated fibers
5.3.5 DCMD results
The fabricated fibers were tested for the DCMD process. Table 5.3 shows the obtained flux as a
function of feed inlet temperature and other operating conditions.
46
Table 5.3 DCMD operating conditions and the obtained flux for the fabricated fibers
During the experiment we observed that the flux increased rapidly as we increased the permeate
inlet pressure to about 8 psig but it remained almost constant for further pressure increase. This
observation is probably due to the fact that as we increased the permeate pressure, a higher
volume of the inner hydrophilic layer occupied by water and the effective water vapor mass
transfer thickness reduced accordingly; but at permeate pressure about 8 psig the inner layer was
probably fully wet and the flux remained constant. Therefore, all the flux measurements were
carried out at permeate flow rates with corresponding pressures of about 8 psig. It can be
observed from Table 5 that water vapor flux as high as 55 kg/m2hr based on the fiber outer
diameter (it is equivalent to 115 kg/m2hr based on the fiber inner diameter) has been achieved at
about 90 C. The reproducibility of data was also tested and confirmed. It is worthy to note that
keeping the permeate side pressure higher than the feed side pressure can insure that the obtained
flux could not be attributed to the leaking of the feed side to the permeate side. A separation
47
factor as high as 99.8% during the experiments confirms this conclusion as well. The separation
factor was calculated based on the following equation [59].
10011001 12
f
pp
f
p
C
C
where Cp is the salt concentration in the permeate side, Cf is the salt concentration in the feed
side, ξp1 and ξp2 are conductivities of permeate stream before and after the test, respectively and ξf
is the conductivity of the feed stream. It is worthy to note that the ionic conductivity of a 3.5 wt%
salt water and deionized water which used for the xpreiments were around 53000 μS/cm and 10
μS/cm respectively. On the other hand, the ionic conductivity of the permeate maintained below
100 μS/cm corresponding to a separation factor of 99.8%.
[85]
[63]
[64]
[86]
[87]
[88]
Table 5.4. Comparison of the maximum flux obtained in this study with the literature for DCMD processes with a hollow fiber configuration **The original data reported in the references [64, 86]
are based on the fibers inner diameter, while we have converted them based on the fibers outer diameter, which is a more realistic and meaningful way to report flux
Fig.5.9 compares the obtained DCMD flux in this study with the maximum flux reported by the
previous investigations on hollow fiber configuration summarized in Table 5.2.
48
[64]
[86]
[63]
[85]
[86] [87]
Fig.5.9 Flux and energy efficiency obtained in the DCMD process using the fabricated fibers and flux comparison with the literature data using hollow fiber membranes
It can be observed that the obtained data in this study are much higher than the most of the
previous reports and there is only one data reported by Li and Sirkar [64] which is quite close to
data obtained in this study. However, as it can be inferred from Table 6, the length of the module
used in [64] is too short (6.4 cm) which actually ignores the limiting effect of permeate
temperature increase along the fibers. This argument is in agreement with the later data reported
by Song et al. [86] using similar conditions but a longer module (25.4cm) that resulted in a lower
flux. The high flux obtained in this study can be mainly attributed to the optimization of different
membrane characteristics such as functional layer thickness, pore size, pore size distribution,
surface and bulk porosities and high contact angle associated with the functional layer without
any advanced module design optimizations. It is believed that if the membrane fabrication
optimizations which were attempted in this work are accompanied with more efficient module
designs to suppress the temperature polarization phenomenon, even higher flux are achievable in
the MD process.
49
50
5.4 Summary
In summary, very different membrane morphologies can be formed in PVDF membranes by
tailoring the strength of coagulation bath. Highly porous membrane structure with three
dimensional networks is formed in PVDF membranes using weak coagulants such as
methanol/water mixtures. Such membrane morphologies possess a very high gas permeability,
leading to a high flux in membrane distillation process. Furthermore, co-extrusion looks to be an
efficient approach to fabricate hydrophilic-hydrophobic hollow fiber membranes to be applied in
MD process. This process provides the freedom to easily tailor membrane properties such as
porosity, pore size distribution and thickness. Finally, high flux and high energy efficiency can be
obtained using the dual layer hydrophilic-hydrophobic fibers.
Chapter 6. Investigation of corrugation phenomenon in the inner contour of hollow
fibers during the non-solvent induced phase-separation process
During the spinning of PVDF and PAN hollow fiber membranes, we came across an instability
that led to the formation of fibers with irregular cross section. In this chapter, we are going to
elaborate the observed phenomenon that here after we call it “corrugation phenomenon”.
6.1 Introduction
As it was discussed in the previous chapters, among different existing geometrical shapes of
membranes, hollow fibers are of great interest because of their high surface area per unit volume
and ease of module fabrication. Comprehensive reviews on hollow fiber membrane fabrication
process can be found in references [89-92]. However, similar to conventional spinning processes,
the fabrication of hollow fibers is inherently accompanied by instabilities that can lead to
technical problems in the production line or undesirable final products that can increase
production costs. Therefore, understanding the nature of these instabilities as well as methods for
controlling them is extremely important. However, these subjects have been often treated as
proprietary know-how in the membrane industry and rarely openly discussed in the membrane
literature.
The major instabilities encountered in fiber spinning are draw-resonance, necking, capillary
break-up and irregular fiber cross-section. The first three phenomena are rather similar in nature
and lead to fiber breakage during the spinning or a non-uniform cross-sectional diameter along
the spun fibers. These instabilities have been investigated theoretically and experimentally by
many researchers and a reasonably good understanding of their mechanism and controlling
methods has been outlined in the literature. They are mainly attributed to the amplification of
51
small fluctuations in polymer jet flow as a result of drawing or capillary forces. Larson, Petrie and
Lipscomb [93-95] have given comprehensive reviews on this subject.
The last type of spinning instability is related to the non-uniform wall thickness and the irregular
cross-section of the as-spun hollow fibers. The severe non-uniformity of the fiber wall thickness
is considered to be major impediment for some membrane applications. For example, in high
pressure-driven membrane processes, the thin regions in the fiber cross-section are considered as
mechanically weak points that are easily collapsed. In addition, the reduced inner cross-sectional
area may lead to high pressure drop in treatment of liquid streams flowing through the fiber
lumen. On the other hand, if a good control of the phenomenon is achieved, it might be possible
to purposely create favorable geometries to the spun fibers that can highly enhance the separation
performance of the membranes. For example, hollow fibers with wavy inner or outer wall could
be preferable because of their high mass transfer area per unit volume. Recently, both Nijdam et
al. [96] and Widjojo and Chung [97] fabricated hollow fibers with wavy geometries at their outer
surface, while Koros et al. [89] and Santoso et al. [98] observed hollow fibers with a wavy inner
skin.
Fibers with wavy geometries might potentially help to increase the effective membrane area and
may induce flow turbulence (in the shell side of the modules) and reduce concentration and
temperature polarizations. Therefore, a more thorough understanding of the governing
mechanisms and approaches to controlling the wavy phenomenon are of great interest. However,
as far as we know, there are very few studies on this subject. Van’t Hof [99] observed the
deformed inner layer in the polyethersulfone (PES) fibers fabricated for gas- separation
applications. He noticed that the addition of solvent or a weaker coagulant to the bore fluid makes
the irregularity more pronounced. Borges et al. [100] observed the same phenomenon in spun
polyetherimide (PEI) and PES fibers. They found that delayed precipitation induced by the
52
addition of a solvent to the bore fluid and increased air-gap distance impeded the deformation of
the fiber cross-section. Both of these studies [99, 100] attributed the instability to the competitive
effects caused by die-swell (which tried to reduce the bore volume) and by bore liquid (which
tried to fill out or expand the bore volume). However, this mechanism does not explain some the
observations reported in the present paper.
In this chapter, we intend to report the results of our experiments on the corrugation phenomenon
to investigate the factors influencing the underlying instability that causes these patterns. We also
propose a novel mechanism that is in good qualitative agreement with our new observations. It is
believed that the proposed mechanism can be potentially applied to mathematically model and
predict the instability.
6.2 Experimental Observations
The following Table summarizes the spinning conditions applied for spinning single layer PAN
hollow fibers. In order to get a better understanding of the phenomenon, we varied the five
spinning parameters namely bore fluid composition, air-gap distance, external coagulant, take-up
speed and dope concentration.
Parameters Value Dope flow rate (ml/min) 2.0
Bore fluid NMP-water (NMP wt. %) 0, 20, 40, 60, 80
Bore flow rate (ml/min) 0.5
Length of air-gap (cm) 0, 3.5, 7, 12, 17, 20
External Coagulant water, Isopropyl alcohol
External coagulant temperature (°C) 25
Spinneret dimensions (mm) 0.85/0.5 (OD/ID)
Spinneret temperature (°C) 25
Take-up speed (ml/min) free fall, 404, 515, 626, 737, 848, 958
Table.6.1 Spinning conditions of hollow fiber membrane fabrication
53
Fig.6.1 shows the cross-section morphology of PAN and PVDF fibers spun through a wet
spinning (no air-gap) process. It is clearly observable that the inner contour of the fibers is
deformed in a way that the wall thickness of the fibers is not uniform circumferentially. In
addition, the cross-sectional view of different sections along the fiber shows that the deformation
had a similar shape all along the fiber. This shows that in contrast to draw-resonance, the
instability is not axially periodic and occurs consistently throughout the spinning process.
Fig.6.1 Irregular shape in the cross section of the PAN 17% (left) and PVDF 20% (right) hollow fibers fabricated through wet spinning, Bore fluid composition of 40% NMP/water
(NMP wt.%), free fall take up rate
The effect of the air-gap distance on macroscopic morphology of the fibers was investigated. The
SEM micrographs in Fig.6.2 show the cross-section of the spun PAN hollow fibers for different
air-gap distances ranging from 0 to 20 cm. It is observed that by increasing the air-gap distance,
the number of corrugations in the inner contour of the fibers decreases from 5 (in the zero air-
gap) to 1 corresponding to a circle (in the 20 cm air-gap).
54
Air gap=17 cm Air gap=20 cm Air gap=12 cm
Air gap=7 cmAir gap=3.5 cmAir gap=1 cmAir gap=0 cm
Fig.6.2 Cross section of spun fibers from 17 wt.% PAN solution, 40 wt.% NMP Bore fluid with different air-gap distances
In another experiment with no air-gap, we observed that by increasing the solvent concentration
in the bore fluid from 0% to 60% the number of corrugations in the inner layer increases from 2
to 8. However, by further increasing the solvent concentration to 80% no deformation is
observable (Fig.6.3).
40% NMP
60% NMP
20% NMP
80% NMP
Pure Water
Fig.6.3 Cross section of wet spun fibers from 17 wt.% PAN solution with different solvent amount in the bore fluid
55
The same trend was observed for higher air-gaps. In addition, we observed that by replacing a
strong external coagulant such as water with a weaker one such as IPA, while keeping other
conditions unchanged, no deformation was observed in the cross section of fabricated fibers. The
SEM micrographs in Fig.6.4 show this behavior.
(A) IPA (B) Water
Fig.6.4 Cross section of wet spun fibers from 17 wt.% PAN solution and 40 wt.% NMP bore fluid using different external coagulants: (A) IPA (B) water
The effect of take-up speed was also investigated. The SEM micrographs in Fig.6.5 show that by
increasing the take-up speed at a constant air-gap distance, the number of waves increased at first
from three to four. However, by further increasing the take-up rate, it decreased so that a circular
cross-section appeared ultimately.
958 cm/min737 cm/min 848 cm/min
626 cm/min 515 cm/min404 cm/minFree Fall
Fig.6.5 Cross section of spun fibers with different take-up rates from 17 wt.% PAN solution, 40 wt.% NMP bore fluid, 7 cm air-gap
56
Finally, we observed that by increasing the dope concentration from 13 wt. % to 17 wt. % and 22
wt. %, the number of waves decreased from 6 to 4 and 1 respectively. The SEM micrographs in
Fig.6.6 show this trend.
Air gap=1 cm 13 wt.% PAN 17 wt.% PAN 25 wt.% PAN
Fig.6.6 Cross section of spun fibers from different concentrations of PAN solution, 40 wt.% NMP aqueous mixture as bore fluid, with 3.5 cm air-gap and free fall
(A) 13 wt.% PAN solution (B) 17 wt.% PAN solution (C) 22 wt.% PAN solution
The same experiments were carried out using a PVDF dope solution and very similar trends but
with different number of waves were observed in these experiments. Fig.7 shows the effect of air-
gap distance on the corrugation of PVDF fibers as an example.
Air gap=7 cmAir gap=7 cm Air gap=0 cm Air gap=3.5 cm
Air gap=20 cmAir gap=17 cm
Fig.6.7 Cross section of spun fibers from 20 wt.% PVDF solution, 40 wt.% NMP Bore fluid with different air-gap distances
57
6.3 Discussion
To propose a suitable mechanism for our observations, we tried to find out in which step of the
fiber-fabrication process the instability arises. The cross-sectional micrographs of as-spun fibers
in Fig.6.8 show that the deformation happens in the fiber-spinning line. Therefore, the
deformation cannot be attributed to the drying or post-treatment stages of the fiber fabrication.
Fig.6.8 Cross section of wet as-spun hollow fibers from 17 wt.% PAN solution and bore fluid containing 40 wt.% NMP
Previous studies [99, 100] postulated that the competitive flow direction between dope solution
and bore fluid in the die-swell region is the responsible for this phenomenon. In order to evaluate
this hypothesis, we took pictures with high magnifications at the die-swell region. Because of the
translucent nature of the PAN solution, we could observe the flow interface between the dope and
the bore fluid in the die-swell region. The pictures in Fig.6.9 show the two-phase flow of the PAN
solution and the bore fluid after discharging from the spinneret. The location of the interface
between the two streams agrees with the spinneret dimensions used in the spinning. Therefore, we
confirm that the interface observed in the pictures is not affected by light scattering. These
pictures show that the interface between the two streams is not corrugated. Therefore, other
factors might be responsible for the instability rather than the competitive force between the dope
and the bore fluid.
58
Bore fluid Rate=0.5 ml/minBore fluid Rate=0 ml/min Bore fluid Rate=2 ml/min
Fig.6.9 Magnified die swell pictures of spinning with 17 wt.% PAN solution, Bore fluid containing 40 wt.% NMP and 7 cm air-gap distance, Different bore flow rates
6.3.1 System description
As polymer dope discharges from the spinneret and contacts with the bore fluid in the die-swell
region, the solvent-nonsolvent exchange between the bore fluid and dope flow leads to the
formation of three phases in the nascent fiber during the dry-jet wet-spinning. Using flow
visualization photos taken from the video microscopy on a flat membrane inversion as an
example, the innermost phase, shown as I1 in Fig.6.10, is formed as a result of vitrification
induced by the bore fluid. The intermediate phase I2 lies between the solidification and
precipitation fronts. In this region, the solution divides into polymer-lean phase and polymer-rich
phases either by nucleation and growth of the polymer-lean phase or by spinodal decomposition.
The third region is between the precipitation and diffusion/convective fronts where macrovoids
propagate and grow [101, 102]. The pictures in Fig.6.10 (A) clearly show the solidification,
precipitation and diffusion/convective fronts and the regions associated with them.
59
O1
I1
O3
I2 I3
O2
Coagulant phase
Precipitation front
Solidification front
I3 or O3 I2 or O2
I1 or O1
(A) (B)
Fig.6.10 (A) Penetration of coagulant into the casting solution after 0.2 second using flow visualization photos taken on flat membrane inversion as an example, (B) Schematic regions in
the extruded nascent fiber
6.3.1.1 Phases I1, I2 or O1, O2
Because of the rapid precipitation of polymer in I1 and I2 (or O1 and O2) phases, they can be
approximated as one pseudo homogeneous elastic cylindrical shell, Is. Therefore, Von Karman
elasticity theory will govern the mechanical behavior of this phase [103].
6.3.1.2. Phase I3 or O3
Since this phase is not fully solidified, we treat it as a viscoelastic fluid with varying density and
rheology that progress with the solvent-nonsolvent exchange between the dope solution and the
bore fluid (or the external coagulant for O3 phase), and the consequent phase separation. For a
detailed description of mass transfer in membrane formation one can refer to the references [104-
106].
60
Since the total density is changing in this phase the following equations of continuity and motion
can be applied. In writing the following equations, we have assumed a convected coordinate
system at the velocity as the interface between the dope and the bore fluid with the positive axial
coordinate (z) in the flow direction.
0.
ut
(6.1)
.... Puut
u
(6.2)
where ρ is the total density, u is the velocity vector ( , , ), t is time, P is the hydrodynamic
pressure and τ is the stress tensor. According to the continuity equation (5.1), velocities are
induced in the r,
ru u zu
and z directions as a result of density change, die swell, non-uniform
progresses of diffusion/convective fronts, and rapid precipitation and densification. Therefore, in
the corresponding equation of motion (5.2), a pressure term has to be introduced in order to
balance the momentum convection and accumulation terms in the left-hand side of this equation.
In a simplified explanation, pressure is generated in this phase mainly as a result of density
change.
6.4 Possible instability mechanisms
As it was mentioned before, our system involves elastic, hydrodynamic, mass transfer and
solidification phenomena. Therefore, the instability could arise from any of these processes. In
the following we analyze the possibility of these mechanisms according to our experimental
observations.
A pure hydrodynamic instability such as those observed in polymeric flows, which mostly initiate
from nonlinear viscoelastic behavior of polymer solutions or viscosity stratifications in two-phase
61
flows is probably not responsible for the observed instability shown in Figs.(6.2 - 6.7). The strong
evidence for this argument is the important effect of the bore fluid composition on the instability.
The results section indicated that different deformations of spun fibers were observed depending
on the solvent amount in the bore fluid, which translates into a different mass-transfer driving
force. On the other hand, it is known that the density and viscosity of the NMP as the solvent are
very close to those of water. Therefore, it seems that the effect of mass transfer, convective flow
and solidification must be taken into account in the stability analysis.
Hence, one possibility is the initiation of an instability that arises from the combined
hydrodynamic, mass-transfer, convective flow and phase-separation processes. In the following,
we describe this possible mechanism qualitatively.
6.4.1 Hypothesis 1. (Mass transfer and hydrodynamic instability)
In this hypothesis, the principal instability is described by the equations of continuity, motion and
mass transfer in phase I3, assuming negligible resistance by the elastic cylindrical shell in phase
Is. This assumption is probably valid if the instability occurs very fast before the rigidity of the
elastic shell in the interface between the bore fluid and the dope becomes so high. In this case the
perturbations occurring in the interface will divide the polymer solution matrix into some regions
with different penetration and contact area with the bore fluid. As shown in Fig.6.11, the regions
associated with the inward movement of the interface have a deeper penetration and increased
contact area with the bore fluid. As a result the solvent-nonsolvent exchange rate between the
dope and bore fluid will be faster and the pressure induced by precipitation/densification in these
regions will be higher compared to the other regions. Subsequently, the generated pressure in
regions associated with inward movement would favor the deformation until a new stable state is
reached.
62
Regions with relatively higher diffusion/convective/solidification rates and higher induced pressure
Perturbed interfaceDeformation progress initiated by the interface perturbations
Fig.6.11 Instability associated with proposed hydrodynamic and mass transfer mechanism
This hypothesis provides a good explanation for the instability initiation. However, it is believed
that without a strong hoop shrinkage force resulting from the solvent-nonsolvent exchange
between the dope flow and the external coagulant, and if the dominant motion and force are in
gravity direction, the initial instabilities may not be able to grow significantly. This argument is in
agreement with our observations in Fig.6.4 and Fig.6.9, which show no corrugation occurrence in
case of a weak external coagulant and in the air-gap region respectively. Therefore, we postulate
that the initial instabilities occurring as a result of densification induced pressure in phase I3 will
be magnified in the external coagulation bath by polymer shrinkage and exponentially growth of
the dope solution density, while a higher viscosity can be considered as a stabilizing factor.
Another possibility is that the elasticity governs the instability, which will be discussed in the
next section as our second hypothesis.
63
6.4.2 Hypothesis 2. (Elastic and Buckling Instability)
In this hypothesis, the principal instability is attributed to the buckling of the elastic cylindrical
shell in phase Is, rather to a hydrodynamic instability. We postulate that the instability occurs in
the external coagulation bath where the inward radial force resulting from the shrinkage of the
dope outer layer and the density-induced pressure in regions I3 and O3 exceeds the critical
buckling pressure of the elastic cylindrical shell in phase Is (defined in 6.3.1.1). Therefore, the
instability can be reduced to the buckling of a thin elastic cylindrical shell in the interface
between the bore fluid and dope solution under a uniform external pressure. This assumption is
probably valid if the resistance by the elastic polymeric shell in phase Is becomes dominant.
Fig.6.12 shows this mechanism schematically.
Elastic inner cylindrical Shell influenced by the bore fluid
Spinneret Dope
Bore fluid
External coagulant
Nascent fiber outer Layer influenced by the external coagulant
Air gap Region
Fig.6.12 Schematic picture showing the exertion of inward radial forces generated by theshrinkage of nascent fiber outer layer in the external coagulation bath
The stability of a long elastic cylinder under external pressure was first studied by Levy [107]. He
reduced the stability problem to an algebraic equation involving elliptic integrals and described
the critical pressure required for buckling by the following equation:
64
3
2* 1R
FqP
(6.3)
2
3
112 v
GhF
(6.4)
In the above equation *P is the buckling critical pressure, B is the flexural rigidity, G is the
Young’s modulus, h is the thickness of phase Is, v is the Poisson’s ratio and k is the buckling
mode or the number of circumferential waves observed in the postbuckling shape. Later, using
elastic theory, Greenhill [108] integrated the post-buckling shapes corresponding to each
equilibrium state. Fig.6.13 shows these shapes corresponding to buckling modes 2 to 5. It is
observable that these shapes agree very well with the cross-section of our spun fibers.
(A)
(B)
Fig.6.13 Close agreement between the spun fibers cross sectional geometry (A) and thepredicted postbuckling shapes of a long elastic cylindrical shell (B) [108]
Fig.6.14 shows the energy of the elastic cylinder as a function of system variables for various
buckling modes as proposed by Tadjbakhsh and Odeh. [109]. This figure shows that as the
pressure increases, the energy of the system increases until it reaches a critical bifurcation point at
which the equilibrium path changes its stability and divides into some stable and unstable
branches.
65
q=4
q=3
q=2
P Rh3 / F
U
Fig.6.14 Energy as a function of the dimensionless group including the external pressure for various modes of deformation [105]
It is worthy to note that the spinneret eccentricity, nonuniform rigidity of the inner precipitated
shell as a result of nonuniform formation of macrovoids and the initial small deformations by the
mass transfer-hydrodynamic instability in the air-gap region may lead to the deviation of the
fibers cross sectional shapes from those theoretically predicted in Fig.6.13 (B).
In the following sections, we explain our observations through the proposed hypotheses.
6.5. Effect of air-gap distance
One possible reason for this trend is that by increasing the air-gap distance, there will be a longer
contact time as well as solvent exchange between the bore fluid and the dope flow before entering
the coagulation bath. As a result, the elastic cylindrical shell in phase Is may become more rigid,
phase I3 may become more viscous and the initial instabilities may be damped out during the air-
gap region at a point in the external coagulation bath where the radial inward shrinkage force is
66
exerted. In addition, for higher air-gaps the cross sectional area of the dope flow decreases before
it enters the coagulation bath. Therefore, the precipitated shell radius, R, in the inner contour of
the fiber decreases accordingly. Each of these factors decreases the tendency for the instability to
occur. The combination of the mentioned factors could be responsible for the decreasing trend in
the corrugation numbers or buckling modes associated with the elastic shell as the air-gap
distance increases. Fig.6.15 illustrates the relationship between P*R3 and the flexural rigidity B as
a function of buckling mode k and shows the stable and unstable regions corresponding to
equation (5.3). The transitional path of the shell equilibrium states as a function of air-gap
distance is shown qualitatively by the symbol of circles in this figure.
q=6
F
Stable
q=2
Rh q=5 q=4
q=3
From low to high P*Rh
3
Bore fluid Composition (NMP/water)
External coagulant (solubility
Air-gap distance
Take-up speed
Fig.6.15 Equilibrium state transitions of the buckled elastic shell as a function of spinning conditions
67
6.6 Effect of bore fluid composition
As the solvent amount in the bore fluid increases, the inner nascent fiber has a lower viscosity and
the flexural rigidity of the inner elastic shell decreases accordingly. Therefore, assuming a
constant external pressure, the corrugation mode is expected to increase. However, in spinning
with a bore fluid containing about 80 wt.% NMP, no nascent solidified inner layer is formed. In
addition, the mass transfer driving force may not be high enough to create the initial instabilities.
As a result corrugation does not happen for this condition and the cross-section maintains a
circular shape.
6.7 Effect of external coagulant
The fact that using a weak external coagulant such as IPA inhibits the instability might be due to
the fact that a weak coagulant induces a slower solvent exchange, diffusion/convection flows,
densification rate and shrinkage which lead to a lower induced external pressure. Therefore, the
induced inward radial force could be lower than the critical force required to buckle the elastic
shell in phase Is or amplify the initial instabilities in phase I3.
6.8 Effect of take-up speed
The increasing and subsequent decreasing trend in corrugation mode as a function of take-up
speed can be explained as follows: A higher take-up rate translates into a shorter contact time and
solvent exchange between the bore fluid and the dope flow in the air-gap region that leads to a
less rigid inner elastic shell. This might be the reason for the initial increase of the buckling.
However, as the take-up rate increases further, the cross-sectional diameter of the dope flow
decreases consequently. Therefore, the rigid elastic shell will have a smaller radius, greater
molecular orientation, and consequently greater resistance to buckling. The smaller radius (i.e.,
smaller R value) and higher molecular orientation (i.e., higher B value) of the elastic shell
resulting from higher take-up rates compensates for the shorter contact time between the bore
68
69
fluid and the dope flow and produces similar and comparable effects that dominate for lower
take-up speeds (i.e., mainly higher B value). The square symbol in Fig.6.15 shows the qualitative
instability path associated with the inner elastic shell as a function of take-up rate.
6.9 Effect of dope concentration
The stabilizing effect of higher concentration or higher viscosity of the spinning solution may be
due to the fact that the rigidity and chain orientation of the inner precipitated shell could be higher
for a concentrated dope compared to a dope with a low concentration. As a result there is less
mass transfer-hydrodynamic instability and the inner shell with a higher rigidity will be more
resistant to buckling.
6.10 Summary
In summary, the non-uniform cross section of hollow fibers can be attributed to an instability
occurring during the fabrication of hollow fibers. It was postulated that the main instability
happens in the coagulation bath where the pressure induced in the nascent fiber outer layer as a
result of diffusion/convection, precipitation, densification and shrinkage will buckle the rigid
elastic shell formed in the interface between the bore fluid and the dope solution. It is also
inferred that since it takes only a fraction of second for a nascent fiber traveling through the air
gap region, the hydrodynamic instability may not have enough time to magnify its effects.
Therefore, it is postulated that the hydrodynamic instability may be the initial stage of causing
instability and the elastic and buckling instability is the magnifying stage for the development of
non-uniform cross section of hollow fibers.
Chapter 7. Conclusion
The following conclusions can be drawn out of this study:
1- Very different membrane morphologies can be formed in PVDF membranes by tailoring
the strength of coagulation bath. Highly porous membrane structure with three
dimensional networks is formed in PVDF membranes using weak coagulants such as
methanol/water mixtures. Such membrane morphologies possess a very high gas
permeability , leading to a high flux in membrane distillation process
2- PVDF surface tension and mechanical properties can be modified by blending the PVDF
polymer with hydrophobic or hydrophilic clay particles. The addition of hydrophobic
particles to the hydrophobic layer further increases the contact angle of the functional
hydrophobic layer by increasing the membrane surface roughness.
3- The surface tension of PVDF polymer can be reduced significantly by blending the
PVDF polymer with hydrophilic PAN as well as the hydrophilic lay particles.
4- Co-extrusion looks to be an efficient approach to fabricate hydrophilic-hydrophobic
hollow fiber membranes to be applied in MD process. This process provides the freedom
to easily tailor membrane properties such as porosity, pore size distribution and thickness.
5- Delamination between the two hydrophilic and hydrophobic layers of the fibers can be
prevented by carefully formulating the dope composition of the two layer in a way that
the two layers are compatible with each other
70
6- High flux and high energy efficiency can be obtained using the dual layer hydrophilic-
hydrophobic fibers.
7- The non-uniform cross section of hollow fibers can be attributed to an instability
occurring during the fabrication of hollow fibers. It was postulated that the main
instability happens in the coagulation bath where the pressure induced in the nascent fiber
outer layer as a result of diffusion/convection, precipitation, densification and shrinkage
will buckle the rigid elastic shell formed in the interface between the bore fluid and the
dope solution. It is also inferred that since it takes only a fraction of second for a nascent
fiber traveling through the air gap region, the hydrodynamic instability may not have
enough time to magnify its effects. Therefore, it is postulated that the hydrodynamic
instability may be the initial stage of causing instability and the elastic and buckling
instability is the magnifying stage for the development of non-uniform cross section of
hollow fibers.
8- The proposed buckling mechanism can effectively explain all our experimental
observations regarding the effect of different spinning parameters on the final shape of
the hollow fibers
9- The proposed qualitative mechanism can be applied to quantitatively model the
corrugation phenomenon in hollow fiber spinning by phase inversion process
In summary, this thesis provides a contribution to tackle one of the major barriers against
commercialization of membrane distillation process regarding the relatively low membrane flux
in this process, by fabricating high flux PVDF composite membranes. Furthermore, this study
provides an explanation for the irregular cross section phenomenon in hollow fiber membranes
fabricated through non-solvent induced phase inversion process.
71
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Appendix
List of papers prepared and published during M. Eng Study
1. S. Bonyadi, T.S. Chung, Flux Enhancement in Membrane Distillation by Fabrication of
Dual Layer Hydrophilic-Hydrophobic Hollow Fiber Membranes, Journal of Membrane
Science, 306 (2007) 134–146
2. S. Bonyadi, T.S. Chung, W.B. Krantz, Investigation of the Corrugation Phenomenon
in the Inner Contour of Hollow Fibers during Non-solvent Induced Phase Separation
Process, Journal of Membrane Science, 299 (2007) 200-210
3. M.M. Teoh, S. Bonyadi, T.S. Chung, J.J. Shieh, Investigation of different module
designs for flux enhancement in the membrane distillation process, Journal of
Membrane Science, 311 (2008) 371–379
4. S. Bonyadi, T.S. Chung, K.Y. Wang, Fabrication of High Performance Dual Layer
Hydrophilic-Hydrophobic Hollow Fibers for Membrane Distillation Process, a US Patent
filed, March 2007
5. S. Bonyadi, T.S. Chung, W.B. Krantz, Investigation of the Corrugation Phenomenon in
the Inner Contour of Hollow Fibers during Non-solvent Induced Phase Separation
AIChE Annual Meeting (2007) Salt Lake City, Utah
6. S. Bonyadi, T.S. Chung, Flux Enhancement in Membrane Distillation by Fabrication of
Dual Layer Hydrophilic-Hydrophobic Hollow Fiber Membranes AIChE Annual Meeting
(2007) Salt Lake City, Utah