membrane distillation for desalination and other separations

128 Recent Patents on Chemical Engineering, 2009, 2, 128-158

1874-4788/09 $100.00+.00 © 2009 Bentham Science Publishers Ltd.

Membrane Distillation for Desalination and Other Separations

Saketa Yarlagadda, Lucy M. Camacho, Veera G. Gude, Zuojun Wei† and Shuguang Deng*

Chemical Engineering Department, New Mexico State University, Las Cruces, NM 88003, USA, †College of Material

Science and Chemical Engineering, Zhejiang University, Hangzhou, P.R. China

Received: March 30, 2009; Accepted: April 27, 2009; Revised: April 27, 2009

Abstract: Membrane distillation is an emerging membrane technology used for desalination of seawater or brackish

water, solution concentration, recovery of volatile compounds from aqueous solutions and other separation and

purification processes. Membrane distillation differs from other membrane technologies in that the driving force for

separation is the difference in vapor pressure of volatile compound across the membrane, rather than total pressure. The

main advantage of membrane distillation over the conventional thermal distillation is that membrane distillation could

occur at a much lower temperature than the conventional thermal distillation. The membranes used in membrane

distillation are hydrophobic, which allow water vapor to pass through but not liquid solution. The vapor pressure gradient

is created by heating the feed solution and cooling/purging the condensate in the permeate side. Therefore, membrane

distillation enables separation to occur below the normal boiling point of the feed solution and could utilize low-grade

heat from alternative energy sources. The objective of this review is to cover the basic principles and configurations of

membrane distillation process, membrane physical characteristics, heat and mass transfer characteristics, and the effect of

operating conditions. Also, major applications of this new technology in desalination, food industry and environmental

protection, and latest patent developments and future trend in membrane distillation are presented.

Keywords: Desalination, membranes, membrane distillation, renewable energy, patents, chemical engineering, heat and mass transfer.

INTRODUCTION

Membrane distillation is relatively a new membrane technology which has been under investigation for more than four decades. The earliest patent discussion of the membrane distillation process was disclosed in U.S. Patents 3340186, 3361645 [1, 2], in which de-mineralized water was produced from saline water using a porous hydrophobic membrane that was not wetted by the saline water. The porous mem-brane separated the hot saline water from the cold de-mineralized water, and a vapor pressure difference was formed across the membrane through which the water vapor passed from the brine side to the de-mineralized water side.

Membrane distillation is being investigated worldwide as a low cost, energy saving alternative to conventional thermal distillation and reverse osmosis process for saline water desalination and other applications. According to Lawson and Lloyd [3], the main advantages of membrane distillation over conventional separation technologies include:

a) 100% theoretical rejection of ions and other non-volatiles from a feed solution;

b) Lower operating temperatures than conventional thermal distillation processes;

c) Lower operating pressures than conventional pressure-driven membrane separation processes;

d) Reduced chemical interaction between membrane and the process solution;

*Address correspondence to this author at the Chemical Engineering Department, Jett Hall, Room 259, New Mexico State University, P.O. Box 30001, MSC 3805, Las Cruces, NM 88003-8001; Tel: (575)646-4346; Fax: (575)646-7706; E-mail: [email protected]

e) Less demanding membrane mechanical property requirements;

f) Reduced vapor space compared to the conventional thermal distillation process;

g) Possibility of utilizing low-grade heat from alternative energy sources.

The above features of membrane distillation process have enabled this technology to be applied for saline water desalination, removal of trace volatile organic compounds from wastewater, concentrated ionic, colloidal and other non-volatile solutions and agro-food solutions. This review will provide an overview of membrane distillation processes with focus on recent developments and applications of this technology disclosed in patents and journal publications. Main operation modes, membrane materials, fundamental concepts of heat and mass transfer characteristics and energy requirements for membrane distillation process are also presented.

1. PRINCIPLE OF MEMBRANE DISTILLATION

In most cases, membrane distillation process is related to the desalination of seawater or brackish water. This tech-nique can be stated as a novel membrane separation process in which two aqueous solutions at different temperatures are separated by a porous hydrophobic membrane barrier. The water vapor flux is thus formed across the membrane from the feed liquid (hot saline water) to the distillate due to the evolved partial vapor pressure [4, 5]. In cases where hydrophobic products are separated from solvent mixtures, such as purification of xylene mixtures [6], the property of the membrane may be changed to hydrophilic to prevent the penetration of hydrophobic liquids. The key principle of

Membrane Distillation: Review Recent Patents on Chemical Engineering, 2009, Vol. 2, No. 2 129

membrane distillation is based on having a membrane non-wetted to the target solvent/ solution and thus allowing only vapor to pass through.

Fig. (1). Principle of membrane distillation process.

The principle of the membrane distillation process is shown in Fig. (1). In general, membrane distillation process constitutes a sequence of seven steps:

1) Movement of the hot feed stream along the surface of microporous membrane, where the temperature and concentration is always considered uniform;

2) Movement of the volatile components of the feed stream to the membrane surface, where temperature and con-centration gradients exist between bulk flow and surface of the membrane;

3) Evaporation of the volatile components in the warm feed water at the membrane interface;

4) Migration of vapor molecules through the non-wetted membrane pore. In addition to mass transfer of the vapor molecules, the heat transfer phenomena comprising latent heat of the vapor, heat conduction, and convection across the membrane also occur;

5) Condensation of the vapor at the membrane surface of condensation side;

6) Movement of the condensed liquid from surface of membrane to the bulk distillate stream, where tempe-rature and concentration gradients exist between the membrane surface and the bulk flow;

7) Movement of the distillate liquid along the microporous membrane, where the temperature and concentration are always considered uniform.

As can be observed from the above steps, the most direct and important parameter that influences the permeate flux is

the target vapor pressure difference across the membrane.

Another important parameter is the energy consumption. The

permeate flux can be expressed as:

N = C(PhbV Pdb

V ) (1)

where N is the vapor permeate flux (kg·m-2·s-1), V

hbP and V

dbP

(kg.m-1.s-2) are vapor partial pressure of target component at

the hot feed stream side and the distillate side, respectively;

and C (s.m-1) is the total mass transfer coefficient. Therefore, any operation that can enhance the vapor pressure difference

will also improve the mass transfer flux across the mem-

brane. Common laboratory or pilot-scale techniques used to

increase the permeate flux include; 1) increasing temperature

difference between the feed stream and distillate side, 2)

mechanical means of creating vacuum at the distillate side to

decrease the local partial vapor pressure, and 3) use of

hollow fiber membranes to increase the specific surface area

of the liquid-gas interface [7-9].

In order to minimize the energy consumption, the heat loss from the hot feed stream side to the distillate side must

be minimized. Such heat loss comes from three kinds of heat

transfer phenomena; (1) the heat conduction across the

membrane due to the temperature difference; (2) the latent

heat transferred by the formed vapor that passes through the

membrane as distillate; and (3) the perceptible heat loss of

the vapor while passing through the microporous membrane,

which is negligible in comparison with the first two terms.

The gross heat loss can be expressed as:

Q =K

(Thb Tdb )+ Nq1 + kh (Thb Tdb ) (2)

where, Q (W m-2) is the gross heat loss from the hot feed stream side to the distillate side, (1/m) is the thickness of

the membrane, K (W m-1 K-1) is the conductive heat transfer

coefficient, N (kg m-2s-1) is the permeate flux of the distillate

liquid, q1 (J kg-1) is the latent heat of distillated vapor, and kh

(W m-2 K-1) is convective heat transfer coefficient.

To reduce the heat losses, lower heat conductive

membrane materials with higher thickness can be employed

[10]. Based on equation (2), smaller temperature difference

between the hot and cold streams or lower feed temperatures

reduce the heat loss. Lower temperature gradient between the hot and cold streams may also result in lower permeate flux

which defeats the purpose of the membrane distillation

process. Therefore, in most cases, as a compromise, criteria

to optimize the membrane distillation process is to obtain a

high permeate flux with minimal energy losses. This rule is

followed in most efforts made for the industrialization of the

membrane distillation process.

It should be noted from Fig. (1) that the driving force of

the permeate flux is the vapor partial pressure difference between the two surfaces of the microporous membrane,

which in turn is equilibrated with the surface temperatures

Thb and Tdb (or boundary layer temperature) of both sides of

the membrane, rather than with the temperatures Th and Td at

both bulk streams. Such phenomenon is called temperature

polarization and needs to be reduced as much as possible

during the membrane distillation process. The value of Thb

should be close to the bulk brine temperature Th in order to

have water vapor pressure at the membrane surface of the

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130 Recent Patents on Chemical Engineering, 2009, Vol. 2, No. 2 Deng et al.

hot brine side as high as possible. Similarly, concentration

polarization is the phenomenon that arises mainly on the hot

feed stream side causing a decrease of the vapor pressure due

to the formation of concentrated salts at the membrane

surface. Contrary to what occurs in the reverse osmosis

process, concentration polarization does not influence

permeate flux adversely in membrane distillation; this is due

to fact that the vapor pressure of the feed is not affected by

small variations in salt concentration [11].

2. TYPES OF MEMBRANE DISTILLATION

CONFIGURATIONS

The membrane distillation process described in the

previous section was viewed as the direct contact membrane

distillation (DCMD) process. In order to obtain an effective

operation mode that could both enhance the permeate flux

and lower the energy consumption, another three types of

membrane distillation processes have been developed [3, 4,

12]. They are the sweep gas membrane distillation (SGMD),

air gap membrane distillation (AGMD) and vacuum mem-

brane distillation (VMD). The four membrane configurations

commonly used in membrane distillation are illustrated in

Fig. (2). Each of these processes has its unique structures and

advantages. Furthermore, a ‘wetted’ membrane distillation

process that uses solvent wetted packed hollow fiber

membrane as the distillate column with different operation

modes has been developed recently.

2.1. Direct Contact Membrane Distillation

Direct contact membrane distillation (DCMD) is the most

commonly used process mode due to its simplicity in

structure, design and process operation. In a DCMD process, seawater or brackish water is passed on one side of a

hydrophobic porous membrane while a colder water stream

flows on the other side to directly condense the permeate

water as shown in Fig. (2a). Since it has a simple structure

that eliminates the need for a separate condenser or a

vacuum equipment or a sweeping gas system, the operation

parameters are only limited to the temperatures, flow rates,

and salt concentrations of the hot brine and the cold water.

The properties of the membrane such as material, porosity

and thickness play an important role as well. Several patents

have covered the process modifications to improve the

permeate flux and energy savings as described below.

Fig. (2). Schematic diagram of different membrane distillation configurations; (a) Direct contact, (b) Air gap, (c) Sweeping gas, and (d)

Vacuum.

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Baig et al. [13] developed a simple DCMD module

which mainly focused on the energy savings during the

process. The authors used bundles of hollow fibers in the

single shell, with each bundle of fibers connect with each

other in a cascade way at two sides of the shell to form only

one inlet and one outlet. The hot solution travelling in the

lumen side thus can have long time to contact with the

permeate vapor in the shell side to recover the latent heat.

The outlet of shell layer can also be operated at vacuum pressure to enhance the driving force of the permeate flux,

which falls into VMD type. However, with the flow in lumen

side going back and forth, the temperature field in the shell

side will be much complex and more difficult to control. As

a result, the efficiency of the module to generate a high

permeate flux and low energy consumption is questionable.

Significant contributions have been made by Sirkar’s

research group on the development of membrane distillation

processes [7, 14-17]. In a US patent 2006076294, Sirkar and

Li [7] presented improvements on DCMD process. They incorporated: (1) a silicone-flouropolymer layer on the

outside surface of Polypropelene membrane to reduce the

water vapor transport resistance and increase the life of the

membrane; (2) larger wall thickness and considerable

porosity to reduce conductive heat loss and water vapor

diffusional resistance through the pores;. (3) hollow fiber

membranes with larger diameter to ensure high flowrate of

the cold distillate and low pressure drop: (4) cross flow of

the brine and distilled water to reduce the temperature

polarization; and (5) asymmetric radial distribution of

hollow fiber by opening different size distribution holes in the two side plates of the fibers. The authors consider that

this design has the ability to provide a uniform flow of the

feed solution through the shell side of the fibers.

Gore et al. [17] patented a spiral wound membrane

module operating in DCMD mode. In their design, the

seawater stream is fed in counter-current direction to the

vaporizing distillate to absorb the heat of the condensation.

A relatively high flow rate of 5.3 kg·m-2·h-1 was achieved

with a temperature difference between the distillate water

and the seawater of 4°C, and energy consumption of only 212 kJ·kg-1 of distillate.

Nano-bubbles have also been introduced to the hot brine

side of a DCMD process using flat membranes [18-20]. Such

nano-bubbles emulated the saline water to improve the mass

transfer of the water vapor across the membrane. The advan-

tages of this design include: (1) reduction of equilibrium

time of water vapor generated from liquid water by

increasing the gas-liquid interface area thus enhancing the

water vapor flux permeating across the hydrophobic mem-brane; (2) increase in turbulence of saline water to reduce the

mass transfer resistance effectively; and (3) sweep out the

foulants through shear forces of the bubble emulsified

solution.

2.2. Air Gap Membrane Distillation

A drawback of DCMD process is a high temperature

polarization due to the limited heat-proof ability of the thin

membrane that separates the hot brine water and the distillate

water. The air gap membrane distillation (AGMD) process

introduces an additional air gap between the membrane and

the condensation surface Fig. (2b). With this, a higher

temperature difference across the membrane and a reduction

in conduction heat losses can be achieved. The main draw-

back of AGMD is the air gap itself as it creates additional

resistance for mass transfer.

Air gap membrane distillation was first described in 1967

in a British patent application [1]. It proposed a counter-

current flow of feed and distillate, and recovery of latent

heat, in addition to the use of an air gap. A German patent

registered in 1982 [21], introduced a 3 mm thickness gap

between a flat porous membrane and a cold condensation

surface for an AGMD application. The gap was filled with

air or with a lighter gas such as hydrogen. Other patents

proposed to feed incoming seawater in counter-current flow

to the vaporizing stream to recover heat [17, 22, 23].

Kurokawa et al. disclosed in a US patent 4879041 [24], an AGMD process suitable for preparation of ultra-pure water

for the semiconductor industry. The effect of the thickness of

the air gap was investigated in the range between 3 and 10

mm using a flat plate membrane. The transport was deter-

mined by diffusion at a thickness of less than 5 mm and by

free convection at a thickness of more than 5 mm. A

maximum distillate flow rate of 3.6 kg·m-2·h-1 was obtained

at a vapor pressure difference of approximately 20 kPa.

A novel AGMD process called Memstill® [17] was developed by TNO, a scientific institution in Netherlands,

which claimed that the process can be industrialized with

capital and operational costs comparable to those of reverse

osmosis [25]. In this process, feed water (seawater, brackish

water, etc.) goes through a non-porous hollow fiber mem-

brane before entering into a porous hollow fiber membrane.

The surface of the non-porous hollow fiber membrane is

subjected to a relatively low temperature and is used as the

condensation surface for the distillate water. The feed water

is placed in counter-current to the porous membrane so that a

large amount of the latent and sensible heat can be recovered

by the feed water, thus saving energy. In order to improve the distillate flow rate, a pressure lower than the atmospheric

pressure and higher than the vapor pressure is maintained in

the air gap by using a vacuum pump. Under these conditions,

a distillate flux of at least 1 kg·m-2·h-1·kPa-1 can be obtained,

and a heat less of 240 kJ·kg-1 of water, or less than 10% of

the latent heat of the feed water can be achieved. The

working temperature is as high as 88 °C, which means that a

temperature resistant membrane material is required.

2.3. Vacuum Membrane Distillation

An effective method to overcome the drawback of mass

transfer resistance of the air gap is to depressurize the con-densate side with a vacuum pump, even though some addi-

tional energy is required to maintain the vacuum Fig. (2d).

This process is known as the vapor membrane distillation

(VMD).


The VMD process described in the patent presented by

Cath et al. [11, 26] (also called vacuum enhanced direct con-

tact membrane distillation), is similar to a DCMD process

except that a vacuum pump depressurizes the shell side

where the distillated water flows. A comprehensive study on

the effect of flow velocity, temperature and concentration of

hot feed brine, absolute pressure at the permeate side, and

membrane type is included in the patent. In this process, an

increase in the flow velocity causes an increase in the per-meate flux. The permeate flux also increases when the

temperature of the hot brine water increases or when the

absolute pressure of the distillate water decreases. Salt

concentration in brine water had little influence on the

distillate water flux. Four types of membrane materials were

evaluated in this study, namely, PS22 (polypropylene, pore

size 0.22 m), TS22 (polytetra-fluoroethylene, pore size 0.22

m), TS45 (Polytetrafluoro-ethylene, pore size 0.45 m),

and TS1.0 (Polytetrafluoro-ethylene, pore size 1.0 m). The

membrane material TS45 is presented as the most favorable

while the membrane material PP22 is presented as the least

favorable for permeate flux. The highest permeate flux reported is about 80 kg·m-2·h-1.

2.4. Sweep Gas Membrane Distillation

In the sweep gas membrane distillation (SGMD) process,

the vapor at the permeate side of the membrane is removed

by a sweep gas and subsequently condensed by an external

condenser. This allows for reduction of water vapor partial

pressure, thus achieving a higher driving force across the

membrane Fig. (2c). SGMD has been used to remove

volatile substances other than water [27]. An advantage of

using a sweeping gas is that the resistance of mass transfer in

the air gap can be substantially reduced. However, dilution of the vapor by the sweep gas generates higher demand on

the condenser capacity.

2.5. ‘Wetted’ Membrane Distillation

In a membrane distillation process, the gas-liquid equili-

brium occurs at the membrane surface of the hot feed stream

side, and the equilibrated gas phase penetrates across the

membrane due to a vapor pressure difference. To prevent

mass transfer resistance between the liquid and vapor phase

and reduction of the permeate flux passing through the

membrane wall, the membrane should not be wetted by the

feed liquids. A new membrane distillation concept was developed by Colling’s research group [28, 29] using a

wetted hollow fiber membrane. The gas-liquid equilibrium

occurs at the membrane surface of the distillate side because

the feed liquid phase penetrates into the wettable membrane.

By carefully selecting the morphology of porous membrane

and controlling the pressure difference across the membrane,

the penetration rate of the liquid phase inside the hollow

fiber is controlled. The liquid phase then forms an uniform

ultra-thin film in the outer side of the membrane and

evaporates to the gas phase thereafter. The vapor diffusion

resistance across the membrane, which is the main difficulty

in the former membrane distillation process, becomes zero in the ‘wetted’ membrane distillation process. The pore size

structure of the hollow fiber membrane used has a denser

structure, with an inner pore size of less than 0.5 m, and an

outer pore size ranging from 0.5 to several hundred of

micrometers. By having a non-symmetrical pore size distri-

bution, the evaporation area of the liquid increases and the

transport resistance of the mass transfer resistance between

the liquid and gas phase decreases, thus leading to an even

higher distillate water flux.

The ‘wetted’ membrane distillation process was first

developed by Colling’s research group [28, 29] as a conven-

tional distillation process to separate light hydrocarbons.

Here, the packing material in a conventional distillation

tower was replaced by a plurality of hollow fibers. The

liquid phase of the mixture went downwards inside the

hollow fiber and the vapor phase went upwards outside. The

mass transfer between the phases occurred in the outer sur-

face of a solvent wettable membrane (polypropylene, poly-

sulfone, polyethylene, polyvinylididene, or polyestersul-

fone).

A comparison has been made between the four types of

membrane distillation process configurations in terms of

effect of process parameters such as Feed and cold water

temperatures, feed and cold water velocities, flux, efficiency,

design considerations and energy requirements per unit

permeate produced. They are presented in Table 1 [30-33].

3. MEMBRANE MATERIALS

In the membrane distillation process, the membranes act

as a physical barrier between the feed and permeate streams and should be sufficiently porous and hydrophobic in nature.

Membranes produced for microfiltration are commonly used

in the MD processs [34, 35]. Membrane material develop-

ment may be the most important strategy to improve the

performance of membrane distillation process. Currently,

hydrophobic polymer membranes such as polypropylene

(PP), polytetraflouroethylene (PTFE), Polyvinylidene-fluo-

ride (PVDF), polycarbonate, and cyclodextrin are available

in capillary or flat-sheet forms. From these, PP membranes

are the most economical. Polypropylene Accurel capillary

membranes have been proven to be successful [34, 35].

PVDF membrane is the only hydrophobic membrane that can be easily dissolved in common organic solvents, but it

can become wetted. It is reported that the contact angle can

be enhanced by delayed demixing-induced through an open

air gelation process. Also, particle-induced surface rough-

ness can be effective by using PVDF/CaCO3 nanocomposite

coating [36]. Hydrophilic materials such as metal fibers,

borosilicate fiberglass, zeolites, aluminosilicate, and sili-

calite, with different pore sizes, thicknesses and porosities

have also been used as membrane materials. Some of the

most common membranes available commercially are listed

in Table 2.

Hydrophobic porous membranes are manufactured by

sintering, stretching, melt spinning and phase inversion

processes [3, 4, 36]. Tarflen membranes are prepared from

commercially available material by extraction of lubricating

compounds, followed by controlled stretching and heating


Table 2. In general, critical requirement for industrial MD

installations is that the membrane should not get wetted

during the separation process [35].

Membranes with porosities between 0.3-0.8 and pore

sizes between 0.1-0.6 m have good performance and are not

supposed to get wetted [5]. Typical membrane thickness range is between 100 and 300 m [37]. The porosity of a PP

hollow fiber can be increased by depositing an ultrathin

highly porous silicone fluoropolymer layer on the outside

surface of the membrane through plasma polymerization [5].

An ideal membrane has high mass transfer rates, low heat

transfer, and is thermally stable and heat resistant. The

choice for a membrane is then a compromise between high

flux and low thermal conductivity [38]. The membrane flux

increases with the pore size and porosity and decreases with

the thickness of the membrane [9]. The flux can be enhanced

with hydrophobic-hydrophilic membranes. The vapor trans-port resistance can also be reduced by using composite

Table 1. Comparison Between the Four Types of Membrane Distillation Configurations

Process Parameter DCMD VMD AGMD SGMD

Feed temperature High feed temperatures

possible

Effective at lower feed

temperatures

High feed temperatures

possible with the air-gap

Optimum feed

temperatures are desirable

due to heat loss in

sweeping

Coolant temperature Optimum coolant temperature,

higher the mass transfer flux,

too low temperatures cause

heat losses [32]

Moderate effect on the mass

flux

Less effect on the mass

flux [32, 33]

Dependent on the coolant

temperature, Independent

of the sweep gas

temperature

Feed inlet velocity Increases mass flux [11,32] and

efficiency

Increase in mass flux as

DCMD, lowers temp and conc

polarizations [30]

Moderate impact on flux

[32]

Cold water inlet

velocity

High impact on flux Moderate impact on flux

Solute concentration Less sensitive compared to

others [30]

Concentration polarization

effect is observed

More sensitive compared

to DCMD [32]

Efficiency Lower than other

configurations

Lower operating pressures give

high efficiency [30]

Efficieny improved due to

air-gap [3]

Maximum permeate

flux (kg/m2-h)

129 [31] 56.2 [42] 28 [31] 21 [31]

Common Applications Concentrating fruit juices,

blood and waste/process water

[3]

Volatile components from

dilute aqueous solutions [30]

Suitable to remove trace

volatile components from

aqueous solutions,

concentration of non-

volatile solutes [3,33]

Used for removing volatile

substances other than

water [3]

Advantages Suitable for high temperature

applications, higher mass flux

than others

Lower conduction heat losses,

higher fluxes than others

Lower conduction heat

losses due to air-gap

Lower mass transfer

resistance

Limitations Large feed flow rates, low

efficiency of heat utilization,

Heat losses due to conduction,

lower cold water temperatures

< 10oC are not suitable

Membrane wetting due to

Pinterface, separation factor can

be affected by lower operating

pressures [30]

Higher mass transfer

resistance than others [3]

External condenser

required to condense the

permeate [3]

Design consideration Reduction in hot feed

temperatures and cooling costs

need to be considered

The vacuum-side pressure is

the major design factor since it

greatly affects the separation

efficiency [30]

Air gap thickness to

maintain flux [33]

Same as DCMD

Velocity and temperature

of the sweep gas, and the

rest same as DCMD [30]

Energy requirements

(kwh/kg)

3.55 kWh/kg of the product 1.1 kWh/kg of the product


membranes prepared from hydrophobic and hydrophilic

polymers.

Composite membranes are prepared by several methods

including coating, grafting, and plasma treatment. Composite

membranes were first patented by Cheng and Wiersma [39,

40]. They modified cellulose acetate and cellulose nitrate

membranes by radiation graft polymerization of styrene onto

the membrane surface, and plasma polymerization of both

vinyltrimethylsilicon/carbon tetrafluoride and octaflouro-

cyclobutane, respectively. Hydrophobic-hydrophilic hollow fiber membranes were prepared by Bonyadi et al. [36] using

co-extrusion spinning process. This process was proposed

for manufacturing composite membranes to obtain optimal

membrane characteristics such as thickness, pore size,

porosity and contact angle. SMM-modified polyetherimide

(PEI) membranes were prepared in a single casting step by

using phase inversion method. SMM-blended hydrophilic

membranes were found promising for desalination using

DCMD process with high fluxes [9]. A multilayer membrane

made of different materials and having two to three layers,

each of them synthesized from different materials, has also been reported [9, 36].

Patel et al. [41] reported a microporous membrane

comprising three layers in a recent patent. The membrane

contains outer layers of relatively small pore sizes, and an

interior layer of large pore size to contact the brine water.

With this structure, the water vapor can penetrate easily into

the membrane, resulting in a larger distillate flux.

4. MEMBRANE MODULES

Membrane distillation has been carried out in a large

variety of module configurations. They include flat sheet,

hollow fiber, tubular, shell & tube, plate and frame, and

capillary membrane modules. The membrane modules are

designed for both laminar and turbulent flow conditions with

longitudinal flow, transversal flow, cross flow and parallel

flow. Figure (3) shows simple structures of flow patterns.

An effective membrane module should have the follo-

wing: uniform flow distribution, high flow rates, low pres-

sure drop, liquid entry pressure (LEP). It should also prevent

high transmembrane hydrostatic pressure [9,41]. Main

aspects to consider when designing a membrane module are

pressure drop along the module and LEP of water [11].

The pressure drop along the length of the channel is

given by

P = f

L

d

u2

2 (3)

where P (Pa) is the pressure drop, L (m) is the channel

length, d (m )is the diameter, and u (m2/sec) is the velocity.

Membrane cells or membrane holders mainly contain inlets

Table 2. Commercially Available Membrane Materials for Membrane Distillation

Manufacturer Trade Name Polymer Pore Size ( m) Thickness ( m) Porosity (%) References

Gelman

Millipore

Membrana, GmbH

Siemens

Sartorious

Gore

Unknown

TF200

TF200

TF1000

TF1000

TF450

TF450

GVHP

PVDF22(GVHP)

PVDF45(HVHP)

GVHP

HVHP

Accurel PP

Accurel PP

PV37

PV65

PV80

HL31

HL50

----

----

----

Tarflen

PTFE

PTFE

PTFE

PTFE

PTFE

PTFE

PVDF

PVDF

PVDF

PVDF

PVDF

PP

PP

PVDF

PVDF

PVDF

Halar

Halar

PTFE

PTFE

PTFE

PTFE

0.2

0.2

1.0

1.0

0.45

0.45

0.11

0.22

0.45

0.22

0.45

0.55

0.60

0.2

0.2

0.2

0.1

0.1

0.2

0.2

0.45

0.8

60

178

60

178

60

178

125

126

116

110

140

150

150

125

170

250

170

150

70

64

77

60

80

80

80

80

80

80

75

62

66

75

75

75

60-80

74

64

60

75

85

70

90

89

50

[50,45,73,83]

[9,85]

[44]

[9,85]

[44]

[9,83,85]

[69]

[50,70]

[50,70]

[9]

[9]

[71]

[49]

[77]

[77]

[77]

[77]

[77]

[50]

[50]

[50]

[74]


and outlets for feed and permeate streams, separators and

membrane itself. A simple membrane holder is shown in

[42,43] Fig. (4).

Fig. (4). Membrane holder [43,44]. entrance (1) and exit (2) for

cold water, entrance (3) and exit (4) for hot water, silicone separator

(5) and (7), membrane (6).

Separators are used in the membrane modules to create channels for the liquid phases. They are placed above and

below the membrane to create the channels for hot and cold

streams [43-45] respectively, and are classified as open

channels and channel spacers as shown in Fig. (5). Open

channel modules are recommended by manufactures for

working temperature less than 50ºC and channel spacers for

temperature lower than 40ºC.

Within the module, the membrane should be supported to

prevent bending, wrinkling and pocketing. A stainless steel perforated plate is used as a membrane support in one

application [46]. The membrane support should have high

mass and heat transfer resistances.

The productivity and process thermal efficiency of

hollow fiber DCMD membrane module has been maximized

using Pareto-solutions [47]. Simulations to investigate the

spatial variations in properties and the sensitivity of the

process performance to operating conditions has also been

Fig. (5). Separators [44,45] (a) Open channels (b) Channel spacers.

studied by the same authors. Song et al. performed scale-up

experiments using large modules for a DCMD process and

observed that the fluxes were close to those obtained with smaller modules. Such behavior was attributed to a rise in

the distillate temperature through the longer hollow fibers.

Two or three channel MD modules were used by Gryta [48]

using net supporters. They concluded that the nets reduce the

transmembrane heat flux and the permeate flux. Higher

water vapor flux was observed when two rectangular

modules were connected in series [49].

Hollow fiber membrane modules are widely used in

many applications due to their high area to volume ratio.

However, they have a non-uniform packing and fiber vibration. Tubular membrane modules have low packing

density and simulate shell and tube heat exchangers. They

are used to concentrate low molecular weight solutions.

Phattaranawik et al. [50] prepared polymethylmethacrylate

membrane modules for both laminar and turbulent flow.

They used the modules in vertical position to prevent free

convection. Improvement of membrane modules is under

investigation. Patents using counter-current flow and cross

flow for feed and distillate streams have been proposed to

recover energy [13, 28]. Patents using nano-bubble at the hot

brine side have also been proposed to enhance mass transfer [18, 51]. In Patents [6, 52], uneven density distributions of

Fig. (3). Flow patterns in membrane modules for membrane distillation (a) Longitudinal-flow (b) Transversal-flow (c) Cross-flow [42].

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hollow fiber was designed to optimize temperature and

concentration distribution. Such variations can, in some

extent, improve the permeate flux and save energy.

5. MEMBRANE PROPERTIES

Membrane properties such as pore size, porosity, tortuo-

sity, and thickness play significant roles in determining the

molar flux through the membrane. The molar flux for a

VMD is given by [30]:

N = Km M (pv pp ) (4)

Km

=2r

3

1

RT

8RT (5)

Where N (mol/m2sec) is the molar permeate water flux,

Km (mol0.5sec/kg0.5 m) is the membrane permeability

coefficient, M (kg/mol) is the molecular weight, pv (Pa ) is

the vapor pressure, pp (Pa) is the permeate pressure, r (m) is

the pore radius, is the porosity, is the tortuosity, (m) is

the membrane thickness, R (J/mol K) is the gas constant, and

T (K) is the temperature.

The flux is related to the membrane properties by

Nr

(6)

the flux increases with increase in pore size and porosity,

and with decrease in thickness and tortuosity of the mem-

brane. A good porous membrane should exhibit high

permeability, low thermal conductivity and high liquid entry pressure of water (LEPw) [9, 36].

The pore size is the most important membrane property

affecting the flow mechanism and therefore the mass transfer

through the membrane. Larger pore sizes increase the mass

transfer rates and increase the molar flux. Larger pore sizes

can, however, lead to wetting of the membrane. In majority

of the applications, water is the main component transported

through the membrane. At 60ºC, its mean free path of water

molecules is approximately 0.3 m [50], therefore an optimal

pore size distribution ranges between 0.3-0.5 m. Figure (6) shows the optimum pore size distribution for the membrane

distillation process as suggested by Bonyadi et al. [36].

The porosity of the membrane is not uniform throughout

the membrane. Both the surface porosity and the bulk

porosity should be considered in designing the membrane.

The flux depends on the mass transfer area, and this in turn

depends on the surface porosity of the membrane. With high

bulk porosities, thermal conductivity of the membrane

decreases and the mass transfer coefficient increases.

Some authors have experimentally characterized the

membranes. Bonyadi et al. [36] used two approaches to

measure the porosity of the membrane. In the first approach,

a mercury porosimeter was used (Autopore III 9420) and in

the second, the membrane was soaked in a 33% LIX54

kerosene solution for one week. Khayet and co-wokers [9]

used gas permeation experiments to measure the pore size,

tortuosity and porosity of membranes.

The tortuosity factor plays an important role in deter-

mining the mass transport mechanism. For a given porosity,

the tortuosity can be measured by [9]:

=

(2 )2

(7)

Fig. (6). Pore size distribution of fabricated fibers [36].

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The tortuosity factor is generally taken to be 2 [53]. The

most successful correlation for estimating the tortuosity is

the Mackie and Meares correlation [54].

5.1. Vapor-Liquid Equilibrium

The flux and composition of the permeating species in

the membrane distillation process is determined by vapor-

liquid equilibrium [46]. The liquid-vapor interface in the

membrane is maintained by the hydrophobic pores and the

selectivity of the process depends on the vapor-liquid equilibrium. For pure components, the vapor pressure is

calculated by the Antonie equation; for binary mixtures, the

partial vapor pressure is given by

p

i= x

i iP

i

0 (8)

Where pi (Pa) is the partial pressure, xi is the liquid mole

fraction, i is the activity coefficient, and Pi0 (Pa) is the

saturation pressure.

The activity coefficient can be calculated by using empi-

rical equations available in literature which are summarized

in Table 3. For organic solution and for finite concentrated

solutions, the Van Laar and NTRL models have been used

successfully [3, 55]. For solutions with no infinite activity

coefficient, the UNIFAC model has fitted the activity

coefficients well. The activity coefficient at infinite dilution

is temperature dependent as given by the equation [58].

lni

TP,n

=H

i

E

RT 2 (9)

lni

|T

= lni

|T

0

+H

i

E

RT 2|T

0

1T

T0

(10)

Where is the activity coefficient at infinite dilution,

HE (J/mol) is the molar excess enthalpy at infinite dilution,

and T0 is taken as 298.15 K.

The activity coefficient is affected by the presence of the

external fields [56].

i

=i,0 i,F

(11)

The values of i,0

and i,F

may be less than or greater

than one based on the positive or negative deviations from

ideality.

The vapor pressure is related to the temperature. For a

small transmembrane temperature difference, the flux can be

calculated by [38]:

N = Cdp

dTT

m

Tf

Tp( ) (12)

where C is the membrane distillation coefficient, and Tf and

Tp (K) are feed and permeate temperatures, respectively.

For dilute solutions that have a small temperature diffe-

rence across the membrane [ T

fT

p( ) = Tm

< 150C ],

the

relationship between vapor pressure and temperature can be

calculated using Clausius-Clapeyron equation [57]:

dp

dTT

m

=p H

vM

RTm

2 (13)

where Hv (J/mol) is the latent heat of vaporization.

Mass transfer across the membrane differs for dilute and

concentrated solutions. Very high feed concentrations lead to

membrane distillation crystallization [38].

For concentrated solutions [38]:

N = Cdp

dTT

fT

p( ) Tth

(1 xm

) (14)

where Tth is the minimum temperature difference and xm is

the mole fraction of the solution.

The minimum temperature difference can be given by

[36, 38, 57]

Tth

=RT 2

M Hv

xf

xp

1 xm

(15)

where, xf and xp are molar fractions of the interfacial feed

and permeate.

For T

fT

p( ) < Tth

, negative driving forces are present

and lead to dilution of the brine solution [38, 57]. The value

of

Tth

is used to estimate the direction of the water vapor

flow in DCMD. The working temperature difference should

be greater than

Tth

in order to have a concentrated solution

[37].

5.2. Membrane Wetting

The hydrophobic nature of the membrane prevents the

penetration of liquid through the pores unless a pressure

higher than the liquid entry pressure (LEP) is applied. The

LEP can be calculated by Laplace-Young equation,

P =2B

L

rmax

cosef

(16)

where B is the geometric factor based on the pore size, L

(N/m) is the surface tension of the liquid, ef is the effective

liquid-solid contact angle, and rmax (m)is the maximum pore

size.

The number of pores penetrated at a given pressure can be

given by [58]:

no. = notot

f (x)dx

r( P)

rmax

(17)

where no is the number of pores penetrated, notot is the total

number of pores, and r ( P) (m) is the radius of the smallest


pore flooded. The integral in the equation gives the fraction

of pores between r and rmax (Fig. (7)).

Water penetration into the membranes can be signi-

ficantly reduced by using membranes with high contact

angle and small pore size. Applied pressure difference

should be less than LEP to prevent membrane wetting. For a

given pore size, the ratio of applied pressure to LEP should

be less than one; if it is greater than one, the liquid will

penetrate into the pores. Surface tension forces in the liquid-

vapor interface prevent the penetration of the liquid through

the pores. The condition for the wettability of the membrane

in terms of surface tension can be given by [37]:

L

=L

p+

Prmax

2B (18)

where Lp (N/m) is the surface tension of the liquid at the

boundary where the liquid penetrates into the pores of the

membrane.

Table 3. Empirical Correlations for Activity Coefficients [14, 55]

Model Empirical Equation

Margules

VanLaar

Wilson

NRTL

G E

RT= ( A

21x

1+ A

12x

2)x

1x

2

ln

1= x

22 A

12+ 2( A

21A

12)x

1 ln

1= A

12

ln

2= x

12 A

21+ 2( A

12A

21)x

2 ln

2= A

21

G E

RT=

A12' A

21' x

1x

2

A12' x

1+ A

21' x

2

ln1

=A

12' A

21' x

1x

2

A12' x

1+ A

21' x

2( )2

ln

1= A

12'

ln2

=A

12' A

21' x

12

A12' x

1+ A

21' x

2( )2

ln

2= A

21'

G E

RT= x

1ln(x

1+

12x

2) x

2ln(x

2+

21x

1)

ln1

= ln(x1+

12x

2) + x

212

x1+

12x

2

21

x2

+21

x1

ln

1= ln

12+1

21

ln2

= ln(x2

+21

x1) + x

121

x2

+21

x1

12

x1+

12x

2

ln

2= ln

21+1

12

ln1

= xk

2k12 G

k1

x1+ x

kG

k1

2

+ 1kG

1k

xk

+ xkG

1k( )2

1k

= a1k

+b

1k

T

ln2

= xk

2k 22 G

k 2

x2

+ xkG

k 2

2

+ 2kG

2k

xk

+ xkG

2k( )2

2k

= a2k

+b

2k

T

G

1k= exp

1k( )

G

2k= exp

2k( )


Fig. (7). Theoretical pore size distribution [58]. The shaded region represents the integral in Eq. (17).

The performance of a membrane distillation process

depends on the extent of the wettability of the membrane

[35]. PVDF membranes are less hydrophobic than PP, PE, or

PTFE membranes. Therefore PVDF membranes have a

greater risk of undesirable wetting [36]. Once the membrane

is wetted it is difficult to restore its original condition by

reducing the hydrostatic pressure [4]. Figure 8 shows various forms of membrane wettability. A deterioration of the

membrane occurs due to partial wetting. Long exposures to

wetting cause decay in the membrane flux.

The hydrophobic properties of the membrane can be

improved to prevent wetting. The contact angle of the flat

sheet PVDF membrane is enhanced by the delayed

demixing-induced through an open air gelation process [59].

Particle-induced surface roughness has been used to enhance

the contact angle, using the ability of PVDF/CaCO3 nano

composite coatings to exhibit hydrophobic behavior [36].

By increasing the length of hollow fiber membrane

module, the risk of membrane wetting increases because of

the hydrostatic pressure drop along the length of the mem-brane module [36]. This risk can be reduced by increasing

the fiber diameter or decreasing the fiber length, maintaining

thus a high membrane area [47]. The deficiencies in hollow

fibers can be overcome with rectangular cross flow of the hot

brine around the porous hydrophobic hollow fiber mem-

branes. This is possible with a thin hydrophobic coating of

silicone-fluoropolymer plasma polymerized on the hot brine

side of the membrane [60].

Clogging of the membrane can occur due to deposition of

insoluble mineral salts on the membrane surface; also due to this reason, the liquid can penetrate easily into the pores.

This phenomenon can be accelerated when the feed contains

salt crystals [34, 61]. Deposits or solids or crystals present in

the feed solution is another reason for membrane wetting. In

this case, wetting can be prevented by using the permeate

from a reverse osmosis unit as the feed solution for

membrane distillation. In doing this, membrane wetting and

scaling can be prevented [34].

5.3. Contact Angle and Surface Tension

The solid-liquid contact angle and surface tension of the

liquid play a vital role in determining the entry pressure difference. To relate the permeate pressure to the pore size

the contact angle must be specified. The contact angle can be

Fig. (8). Various forms of membrane wettability in MD process [35] (A) Non-wetted (B) Surface-wetted (C) Partial-wetted (D) Wetted.

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calculated using Young’s equation

Lcos =

S SL (19)

where S (N/m) is the surface tension of the solid, SL (N/m)

is the surface tension of the solid-liquid, and L (N/m) is the

surface tension of the liquid.

The contact angle for low energy surfaces such as

hydrophobic surfaces is given by [58]:

Lcos =

L+

L

0 (20)

where L0 (N/m)is the surface tension of spontaneous

wetting.

The contact angle for the rough surface is obtained from

the relation

cos *

= f1

cos f2

(21)

where * is the contact angle of the rough surface, is the

contact angle of the smooth surface, and f1 and f2 are the

fractions of liquid-solid and liquid-air surfaces, respectively.

The concept of the penetration surface tension in addition

to the surface tension of liquid provided satisfying results for the wettability [4]. The penetration surface tension is

measured using the penetration drop method.

Different instruments have been used to measure the

contact angle in membrane materials. Chung et al. [36]

measured the contact angle for PVDF membranes in hollow

fibers using a tensiometer and in flat-sheet membranes using

a contact angle goniometer. They obtained a contact angle of

136o for the hollow fibers and 140o for the flat sheets. They

concluded that high contact angles may be due to high

roughness on the membrane surface caused by the formation of a porous three-dimensional fiber-like network with some

degrees of spherulitical structure and the presence of clay

particles in membrane matrix. Song et al. [61] used Cahn

DCA315 balance to measure the dry surface contact angle

and the wet surface contact angle of water on hollow fiber

membrane surfaces. The effect of temperature, source and

concentration of contaminants, and thermal cycle on the

hydrophobicity of membrane surface were also studied [61].

5.4. Membrane Fouling

Fouling of a membrane is caused by the irreversible

attachment of the solute layer to the membrane surface. The types of fouling in a membrane process include inorganic

fouling or scaling, organic fouling, particulate and colloidal

fouling, and biological fouling (biofouling) [62]. Scaling or

inorganic fouling is a major problem in seawater desali-

nation. Scaling occurs when hard scales of CaCO3 or CaSO4

are formed. In MD scaling occurs when the process is used

for treating effluents from ion-exchange processes con-

taining iron oxide. Scaling can be prevented by rinsing the

module with concentrated acid solution i.e., acidification

[61], pretreating the waste water with lime precipitation

followed by sedimentation and filtration [63], using low

porosity surfaces [35], or adding thermal softeners and by

using net filters [64].

Organic fouling occurs due to adsorption or desorption of

dissolved or colloidal organic matter on the surface of the

membrane. This may occur at the molecular level or on the

surface [62]. Formation of protein deposits can be prevented

by boiling the feed solution followed by filtration [62]. Particulate and colloidal fouling occur due to the presence of

particulate species in the feed solution. These particles or

colloids get trapped in the membrane-liquid interface by the

surface tension. Prefiltration of the feed solution is sufficient

to reduce the effect of the particle and colloidal fouling [4].

Biofouling arises due to the presence of biological matter in

the feed solutions. Chemical disinfection combined with UV

treatment can control biofouling [4]. In general, deposits

formed on membrane surfaces can be eliminated by using

integrated RO-MD unit [65], chemical pretreatment of the

feed solution, and clarification followed by filtration.

Fouling layers can be of two types, namely porous and

non-porous. Porous fouling layer causes an increase in the

thermal resistance; non-porous layers cause an increase in

the mass transfer resistance [62]. In both cases a flux decay

occurs. The presence of fouling layers in the surface of a

membrane causes the wetting of the membrane. Gryta et al.

used scanning electron microscope (SEM) coupled with

energy dispersion spectrometry (EDS) to study the compo-

sition and morphology of the deposited layer on the mem-

brane surface. They also used attenuated total reflec-tance fourier transform infrared (ATR-FTIR) spectrometry to

identify the functional groups associated with the membrane

foulant.

6. HEAT AND MASS TRANSFER AND ENERGY

EFFICIENCY

6.1. Heat Transfer in Membrane Distillation

Heat transfer in membrane distillation occurs due to the

diffusion of species and conduction of heat through tempe-

rature differences [66]. Heat transfer in membrane distillation

occurs in three forms [51, 66-68]

1) In the feed side, the heat transfer rate Qf is the

summation of convection in the feed (hot fluid)

boundary layer, Qf,conv, and mass transfer across the feed

thermal boundary layer, Qf, M.T,

Q

f= Q

f ,conv.+ Q

f , M .T .= h

f(T

fT

mf) + N H

L, f (22)

2) The heat transfer rate across the membrane Qm is due to

conduction through the membrane material, Qm,cond , and

heat of evaporation of the volatile stream, Qm,M.T., which

is also equal to the heat of conduction through the

membrane pores, and latent heat of vaporization, Qm,V,

Q

m= Q

m,cond .+ Q

m, M .T .= h

m(T

mfT

mp) + N H

v (23)

3) In the permeate side: due to convection in the permeate

(cold fluid) boundary layer, Qp,conv, and due to mass


transfer across the permeate thermal boundary layer,

Qp,M.T,

Q

p= Q

p,conv.+ Q

p, M .T .= h

p(T

mpT

p) + N H

L, p (24)

where hf, hm and hp (W/m2 K) are the heat transfer

coefficients of the feed, membrane, and permeate side,

respectively; Tf, Tmf, Tmp, and Tp (K) are the temperatures of

feed, feed-membrane interface, membrane-permeate inter-

face, and permeate streams, respectively; and HL,f and HL,P

(J/kg) are the enthalpies of feed and permeate solutions,

respectively, all of them are calculated at the average tempe-

rature of the feed and the permeate. The resistance to heat transfer comes from the thermal boundary layers, conduction

across the membrane, and latent heat of vaporization as

shown in Fig. (9) [4].

Temperature polarization is the main reason for heat

transfer through the boundary layer [42]. The heat transfer

across the boundary layer is a rate limiting step for the mass

transfer because a large quantity of heat is required to

vaporize the liquid at the vapor-liquid interface [69]. Heat

transfer with simultaneous mass transfer takes place in the

membrane distillation process resulting in complex heat transfer mechanism. Different heat transfer models have

been proposed to explain the mechanism of heat transfer and

to facilitate the evaluation of the membrane surface tem-

perature [50, 68]. The models assume linear temperature

profile and isenthalpic flow. Phattaranawik et al. [50] pro-

posed a heat transfer model to identify the influence of mass

transfer on the heat transfer rate and the significance of the

heat transfer mechanism. At steady state conditions, the heat

flux in each of the regions of the membrane is identical, and

the temperatures at the membrane surface can be evaluated

as [68].

Q = Q

f= Q

m= Q

p (25)

In the feed and permeate regions of the membrane the heat transfer due to mass transport is very small compared to

the heat transfer due to convection, therefore the above

equations reduce to

Q = Q

f= Q

f ,conv.= h

f(T

fT

mf) (26)

Q = Q

m= Q

m,cond .+ Q

m, M .T .= h

m(T

mfT

mp) + N H

v (27)

Q = Q

p= Q

p,conv.= h

p(T

mpT

p) (28)

Q = hf (Tf Tmf ) = hm (Tmf Tmp )+ N Hv = hp (Tmp Tp )

(29)

Rearranging the equations, the equations for the

membrane surface temperatures can be obtained as:

Tmf

=h

m(T

p+ (h

f/ h

p)T

f) + h

fT

fN H

v

hm

+ hf(1+ h

m/ h

p)

(30)

Tmp

=h

m(T

f+ (h

p/ h

f)T

p) + h

pT

p+ N H

v

hm

+ hp(1+ h

m/ h

f)

(31)

The heat transfer coefficient for the membrane can be obtained from the thermal conductivity value

h

m=

km =

kg

+ ks(1 )

(32)

where km (W/m K) is the thermal conductivity of the

membrane, kg and ks (W/m K) are the thermal conductivities

of the hydrophobic membrane material and the air trapped

inside the membrane pores, respectively. The heat flux can be obtained by the relation

Fig. (9). Heat transfer resistances in membrane distillation [4].

�12 3��412

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Q =1

hf

+1

hm

+N H

v

Tmf

Tmp

+1

hp

1

(Tf

Tp) (33)

The boundary layer heat transfer coefficients are gene-

rally evaluated using the empirical correlations for the

determination of the Nusselt number. Selection of appro-

priate correlation is important to explain the heat transport in

the membrane module. The general correlation for the

Nusselt number is given by [45, 70]:

Nu = A Re Pr (34)

where Nu is Nusselt number

Nu =h

id

km

, Re is Reynolds

number

Re =dv

μ, Pr is Prandtl number

Pr =C

pμ

km

, and A,

, and are parameters for the working regime and

correlation used. Table 4 [13, 47, 50, 68, 69, 71-75] summarizes some correlations for heat transfer coefficients.

6.2. Mass Transfer in Membrane Distillation Process

Mass transfer in membrane distillation occurs mainly due

to diffusion and convection caused by a concentration diffe-

rence in the vapor [38, 76]. The driving force for mass

transfer in membrane distillation is the vapor pressure diffe-

rence across the membrane. The mass transfer in membrane

distillation can be categorized into three regions [66].

1) Mass transfer due to diffusion from the feed stream to the membrane interface.

2) Mass transfer due to diffusion and convective vapor

transport through the pores of the membrane.

3) Mass transfer due to vapor condensation of the vapor

from the membrane interface to the permeate side.

The membrane distillation process is defined based on

ideal conditions; therefore the viscous flow of the liquid and

the thermodiffusion within the pores are excluded. The

transport of solvent is given by diffusion of the permeate

material because of a vapor pressure difference across the membrane. The mass transfer resistances in membrane

distillation are presented in Fig. (10). The mechanism of

diffusion in the membrane distillation process is explained

by the Fick’s laws. The molar flux can be described by the

equation [37]:

N =

cD

1 x

dx

dz (35)

where c (mol/m3) is the molar concentration of solution, D

(m2/sec) is the diffusion coefficient, x is the mole fraction of water vapor, and z is a constant. This equation is only valid

for dilute solutions where activity coefficients are unity [37].

Main problems in the solution of the above equation are

that the diffusion coefficient varies within the membrane

pores and within the free region, and also diffusion coeffi-

cient depends on the temperature and vapor concentration.

A linear relationship between the water vapor flux and the vapor pressure difference can be obtained as

N = C( p

mfp

mp) (36)

where C (mol/m2 sec Pa) is the membrane distillation coeffi-

cient, pmf and pmp (Pa) are the partial pressures of feed and

permeate at the feed-membrane and membrane-permeate

interface, respectively.

The partial pressure can be evaluated using the Antonie

equation [77]:

Fig. (10). Mass transfer resistances in membrane distillation [4].

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=23 ��28��& :15��35688��&8��4��62��

;132 68<56<�8

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Table 4. Correlations for Heat Transfer Coefficients

Equation Comment References

Nu =1.86 Re Prd

h

L

1/3μ

b

μm

0.14

Nu = 0.023Re0.8 Pr0.3

Nu = 0.5 Re Prd

h

L

μb

μw

0.14

Nu = 0.116 Re0.66 125( )Pr0.33 1+d

h

L

0.66

μ

μw

0.14

Nu = 0.15Re0.88 Pr0.33

Nu = 1.62 Re Prd

h

L

0.33

Nu = 0.023 Re( )

0.8Pr( )

0.33

Nu = 1.61Re0.35 r

b

1/3

Pr1/3

Nu = 0.298 Re0.646 Pr0.316

Nu = 0.027 Re0.8 Prc μb

μm

0.14

Nu = 0.04 Re0.75 Pr0.33

Nuf

= 1.04 Re0.4 Pr0.36 Pr

Prw

0.25

Fc

Nuf

= 0.71Re0.5 Pr0.36 Pr

Prw

0.25

Fc

Nus= 0.664k

dc(Res )0.5 Pr0.33 2d

h

s

lm

kdc

= 1.654d

f

H

0.039

0.75 sin2

0.086

Laminar

Turbulent

Re Pr

dh

L> 10

2000 < Re < 10,000

Helically wound capillary modules

Re > 2100

Re > 2100

500 < Re < 5000

150 < Re < 3500

Turbulent

c = 0.4 for heating

c = 0.3 for cooling

Turbulent

Due to fouling

Re < 40

Re > 40

For spacer-filled channels

[13,71,72]

[72]

[71]

[71]

[71]

[47]

[47]

[73]

[74]

[68]

[69]

[13,5]

[5]

[50,75]


Fig. (11). Relation between temperature and vapor pressure by Antonie equation [78].

pv= exp A

B

T + C (37)

The values of A, B and C in the equation are constants.

For water, the values are A = 23.1964, B = 3816.44, C = -46.13. Figure 11 gives the relationship between tempe-rature

and water vapor pressure based on Antonie equation.

The molar vapor diffusion flux of water is given by the

general Maxwell’s diffusion equation as [72]:

N =DPM

RTln

P pp

P pf

(38)

The mass transfer in the permeate boundary layer is not

considered because the mole fraction of the transporting

species in the permeate stream is approximately equal to one.

In hollow fiber membrane modules the mass transfer is

explained by the overall mass transfer coefficient [13]. In the

membranes used for the VMD, the pore size is usually

smaller than the mean free path. The prevailing mechanism for the mass transfer through the membrane pores at the low

vacuum pressure is the Knudsen diffusion [67,78].

In air gap membrane distillation, water vapor diffuses

only through the membrane, air acts as a stagnant film and

the vapor diffusion flux is given by the Stefan diffusion [76].

N =DP

RTa

( pv

pf)

pln

(39)

where a (m) is the air gap thickness, and pln is the log mean

partial pressure difference of the stagnant air. Total

resistance to mass transfer in AGMD is given by the sum of

resistances caused by the hot solution, membrane, air/vapor

gap, condensate film, cooling plate and cold solution [79].

The resistance to mass transfer through the membrane comes

from the air that is trapped in the porous membrane structure

i.e., pore size, porosity and tortuosity [13].

Diffusion of gaseous species through the porous mem-

brane depends on the properties of the membrane namely,

pore structure, thickness of membrane, void pore area, and

solid membrane area perpendicular to the direction of

diffusion [76]. Diffusion transport of vapors in porous media

can be explained by various mass transfer models. In Dusty

Gas Model (DGM) mass transfer across the membrane can

be explained by Knudsen diffusion, molecular diffusion,

surface diffusion, and viscous flow. Surface diffusion is always neglected in membrane distillation. Mass transfer in

boundary layer is analyzed by the film theory; mass transfer

across the membrane by the Dusty Gas Model [50,66,70].

The selection of a model is based on the properties of the

membrane and volatile vapors. Knudsen number, Kn

provides the guidelines for determining the mechanism for

the vapor transport in the porous membrane medium. Kn is

defined as the ratio of the mean free path ( ) of the vapor

molecules transported to the pore size (d)of the membrane

[66]:

Kn = /d (40)

Knudsen diffusion occurs when collisions of molecules

with the inside wall are predominant as compared to the

collisions within the molecules. Molecular diffusion model is

suitable for systems in which molecule-molecule collisions

dominate over the molecule-wall collisions. Viscous flow

takes place when a portion of gas flows in a continuum

region due to pressure gradient.

When Kn < 0.1, the transport is affected by the molecular

diffusion and the molar flux is given by [66]:

N ' =PD

RT

( pmf

pmp

)

| pa

|ln

(41)

When Kn > 10, the transport is effected by the Knudsen

diffusion and the molar flux is given by [66]:

!%$

*%

'$

(%

)$

#%

!$

%% !% "% #% )% $% (% '% +% *% !%%

�� .�/�

��

��

��

.>��/

$# ��

!( ��

!% ��

' ��

( ��

$ ��


N ' =1

RT

2 r

3

8RT

M

1/ 2( p

mfp

mp)

(42)

In the transition region where 0.01< Kn < 1, molecules of

water collide with each other and diffuse among the air

pores. Mass transport in transition region takes place by

Knudsen-molecular diffusion mechanism. The flux can be

given by [66]:

N ' =3

2 r

RT

8M

1/ 2

+p

a

PD

RT

M

1

( pmf

pmp

) (43)

The flux in this region is directly related to the pressure

difference as [37].

N = a b p (44)

where a and b are empirical coefficients from the

experimental data.

For Knudsen diffusion (lighter molecules), b = 0

For molecular diffusion, b = 1

Mass transfer coefficients are calculated using Sherwood

number. The general correlation for Sherwood number is

given by [4]:

Sh = A Re Sc (45)

where Sh is the Sherwood number Sh =

km

d

D and Sc is

Schmidt number

Sc =μ

D, A, , are parameters that

depend on the working regime and correlation used. Some

correlations for mass transfer coefficients are presented in

Table 5.

Mass transfer in hollow fiber membrane module is

described by the overall mass transfer coefficient [13]. The

majority of the aerated systems are mass transfer limited. The flux can be improved by deaerating the system because

deaerating the membrane decreases the molecular diffusion

resistance which makes the Knudsen flow dominant [38].

Srisurichan et al. observed that the Knusden-diffusion model

best fits for laminar flow regions and molecular-diffusion

model fits for the turbulent region [66]. Ding et al. developed

Knudsen-molecular diffusion-Poiseuille (KMPT) model to

describe the mass transfer through the membrane [80].

Lawson and Lloyd solved the vacuum membrane distillation

numerically by Micromath scientific, membrane parameters

were evaluated by gas permeation experiments and physical properties of the feed and permeate were calculated from the

empirical correlations [81].

6.3. Spacers

The use of separators in the membrane distillation

process increases the flux. Screen separators or spacers are

Table 5. Correlations for Mass Transfer Coefficients

Equation Comment Reference

Sh = 1.61Re0.35 r

b

1/3

Sc1/3

Sh = 3.56 Re0.265 Sc1/3

Sh = 1.86 Re Scd

h

L

1/3

Sh = 0.023Re0.8 Sc0.3

Sh = 0.04 Re0.75 Sc0.33

Shf

= 1.04 Re0.4 Sc0.36 Sc

Scw

0.25

Fc

Sh = 0.664 Re0.5 Sc0.33 dh

L

0.5

500 < Re < 5000

250 < Re < 1200

Laminar

Turbulent

Turbulent

Due to fouling and scaling Re < 40

Laminar

[73]

[73]

[72]

[72]

[69]

[49]

[75]


placed in flow channels and are made of polypropylene.

They act as mixing devices to promote the heat transfer [44].

They enhance the momentum, and heat and mass transfer by

creating eddy currents in laminar regions [50]. Phattaranawik

et al. [11] achieved 60% of flux enhancement by using net-

type channel spacers. A flux increase of 26-56% was

observed in laminar regions using spacers [50]. The flow

patterns in spacer filled channels are presented in Fig. (12)

[50].

Martinez-Diez et al. [44] used fine and coarse screens to

study their effect on the MD process. Characteristics of the

spacers used in their study are presented in Table 6. They

observed that the turbulence created by the spacer helps to

reduce the effects of temperature polarization. Flux enhance-

ment was higher for coarse spacers than for fine ones.

6.4. Temperature and Concentration Polarization

6.4.1. Temperature Polarization

The membrane/interface temperatures are different from

the bulk temperatures of the feed and permeate side. This is the main phenomenon for the temperature polarization to

take place. Temperature polarization (TP) across the mem-

brane is a major energetic inefficiency in the membrane

distillation process [11]. The liquid boundary layers adjacent

to the membranes are the main reasons for TP because a part

of the externally imposed temperature difference is dissi-

pated through the boundary layers. The extent of temperature

polarization can be understood by the temperature polari-

zation coefficient, .

The temperature polarization coefficient (TPC) can be

defined as the ratio of the actual and theoretical driving force

for heat transfer [66].

=T

mfT

mp

Tf

Tp

(46)

Where is the temperature polarization coefficient, Tmf is

membrane-feed interface temperature, Tmp is membrane-

permeate interface temperature, Tf and Tp are bulk feed and

permeate temperatures respectively.

The membrane/interface temperatures are evaluated by

performing heat balance related to the bulk temperatures.

Fig. (12). Flow direction in spacer-filled channels and spacer fabrication [50].

�

��

�

��

��

�� !�

"��

#��

��

�� #��

"��

$��

�� "��


Temperature polarization occurs in both the feed and

permeate streams and reduce the mass transfer rates because

of considerable losses in the thermal driving forces [75]. The

extent of the temperature polarization depends on the pro-perties of the membrane and on the heat transfer conditions

in the liquid streams. The heat transfer conditions can be

determined by the hydrodynamic conditions governing the

liquid flow within the channels of the membrane modules

[38].

TPC indicates the thermal efficiency of the system, but

does not measure directly the reduction in the driving force.

Gryta et al. observed that for same values of , the driving

force for the MD process may have different values [62]. For

the VMD process, is almost independent of the feed tem-perature, but increases as the permeability of the membrane

increases [81]. The temperature polarization tends to reduce

the scaling potential at the solution-membrane surface for the

case of inversely soluble salts (CaSO4, CaCO3) [49]. The

presence of fouling layer decreases the TPC. Temperature

polarization contributes to the super-saturation build-up at

the membrane surface and a maximum can be observed at

the membrane wall.

The models developed for the MD process of pure liquids or low molecular weight solutions can be used to measure

TPC. Agashichev et al. [82] developed a model based on the

mass and energy balance equation for the hydrodynamic,

temperature, and concentration of boundary layers in a plate-

and-frame unit, taking into consideration temperature-

concentration polarizations. The model predicts the tempe-

rature and concentration distribution along the membrane

surface. It also permits simulation and analysis of the

influence of various factors on the permeate flux.

The temperature polarization is controlled by the overall heat transfer coefficient (U) of the boundary layers, which

strongly depends on the tangential velocity of the fluid in

contact with the membrane. Schofield et al. developed a

model to evaluate the U value for the MD process, assuming

that the coefficient is equal on both sides of the membrane

and independent of the mean temperature [83].

For a given value of the tangential velocity, the trans-

membrane flux is given as

(Tf

Tp)

N Hv

=U + 2k

m

HvCU

1

dp

dT

+2

U (47)

Where Tf and Tp are the bulk temperatures of feed and

permeate side (K),N is molar flux (mol/m2 sec), is the

membrane thickness (m), km is the thermal conductivity of

the membrane (W/m sec), Hv is the latent heat of vapori-

zation (J/mol), p is the pressure (Pa), T is the temperature

(K), C is the membrane distillation coefficient (mol/m2 sec

Pa), U is the overall heat transfer coefficient (W/m2 sec)

From the intercept and slope of the Schofield plots, the

values of U and C can be calculated. Based on his model, the

TPC can be calculated as

TPC =

12 H

v

Ucalc

N

(Tf

Tp)

1+2k

m

Ucalc

(48)

where Hv is the latent heat of vaporization (J/mol), Ucalc is

the calculated overall heat transfer coefficient (W/m2 sec), N

is the molar flux (mol/m2 sec), Tf and Tp are the bulk

temperatures of feed and permeate side (K), km is the thermal

conductivity of the membrane (W/m sec), is the membrane

thickness (m).

TPC values range between 0 and 1. As 0, with Tmf << Tf, the effect of the temperature polarization is more

significant and the process rate is completely controlled by

the heat transfer resistance in the liquid phase. As 1, with

Tmf Tf, the effect of the temperature polarization is not

significant and the process rate is completely controlled by

the transmembrane mass transfer resistance. High TPC

values are due to high heat transfer coefficients in the

boundary layer which leads to high turbulence [1]. Within

the range of 0 and 1, it is not clear whether the mass flux is

controlled by the mass transfer resistance through the

membrane or by the heat transfer resistance of the liquid

phase [76, 84].

Burgoyne et al. [38] and Bryk et al. [37] reported highest

TPC values for hollow fiber and flat sheet modules in

membrane distillation as follows;

• 1 mm tube, Re 5000 in the tube

• 0.3 mm tube, Re 300 in the tube

• mm film with laminar flow

Srisurichan et al. observed that the TPC ranged between

0.4-0.53 for low crossflow velocities (laminar region) and

Table 6. Geometric Characteristics of Spacers Used by Martinez’-Diez [43]

Channel Dimensions (mm) Spacer Spacer Thickness

(mm)

Strand Diameter

(mm)

Voidage

(Non-Dimensional ) Height Length Width

Fine

Coarse

0.40

0.50

0.20

0.25

0.65

0.73

0.60 97 61

0.60 97 61


0.87-0.92 for high crossflow velocities (turbulent region)

[69]. TPC values smaller than 0.2 indicate that the process is

heat transfer limited and the module has a poor design; and if

the values are greater than 0.6, the process is mass transfer

limited due to low membrane permeability [70].

6.4.2. Concentration Polarization

Similar to the variations in the temperature, the concen-tration of the solute at the membrane/interface is different

from the bulk feed side concentration leading to concen-

tration polarization. Concentration polarization originates

from the presence of concentration boundary layer at the

feed membrane interface.

The effect of concentration polarization can be analyzed

through the concentration polarization coefficient, . The

coefficient is defined as the ratio of the concentration at the

membrane surface to the concentration in the bulk feed

stream.

=C

m

Cb

(49)

where Cb (mol/L) is the bulk feed concentration, Cm (mol/L)

is the membrane interface temperature.

One of the main advantages of the MD process is the

complete rejection of the solute. The retained non-volatile

solutes are accumulated at the membrane surface. Such a

concentration build-up may partially or completely block the

membrane pores resulting in lower permeate flux [67]. This

concentration build-up also leads to the diffusive flow back

to the feed side [83]. In practical applications, the retention of

non-volatile solute never attains 100%. Therefore, it is

necessary to take into account this fact in theoretical analysis

[37].

The interfacial concentration can be calculated from the

relation [85, 86].

Cm

= Cb

expN

km

(50)

Under steady state conditions, the convective solute

transport to the membrane surface can be balanced by the

solute flux through the membrane and back diffusive flow from the membrane surface to the bulk of the feed [46]. As

1, with Cm Cb, the effect of the concentration polari-

zation becomes insignificant and the interface concentration

becomes close to the bulk concentration [67] . For > 1, with

Cm > Cb, the effect of the concentration polarization becomes

significant and the interface concentration is greater than the

bulk concentration [67].

Izquierdo-Gil et al. [69] proposed a method to evaluate

the concentration polarization effect from the fit of the

experimental data. Liquid-vapor equilibrium concentration and membrane concentration were evaluated and compared

with the experimental permeate concentrations data. In

practical cases, the temperature polarization is paid attention,

however, when the feed concentration increases the effect of

the concentration polarization has to be added. This reduces

the driving force and thereby the mass flux [87]. The process

behavior can be studied using polarization factors [79].

Since heat and mass transfer take place simultaneously,

both the temperature and concentration polarization effects

should be considered. Temperature and concentration polari-

zations influence the characteristics of the membrane. By

altering the nature of solution and permeate material fluxes, polarization can be partially eliminated [38]. Improving the

flow characteristics can reduce the effects of polarization.

High flow rates or turbulent conditions can decrease the

polarization [50,75]. Recently turbulence was created by

using spacer filled channels. The values of ‘ ’ and ‘ ’ are

dependent on the film transfer coefficients ‘hf’, ‘hp’ and ‘km’.

Some authors used the vapor pressure polarization

coefficient, ‘f’, to measure the reduction in the effective

driving force [45, 87, 88]. The vapor pressure polarization

coefficient is given by the equation [87]:

f =p

mfp

mp

pf

pp

(51)

where f is the vapor pressure polarization coefficient, pmf and

pmp are the vapor pressures of membrane-feed and mem-

brane-permeate interface (Pa), pf and pp are the vapor pressures of feed and permeate (Pa).

6.5. Evaporation Efficiency

Maximum molar flux is the main criterion for the MD

process. Optimization of the process should also involve the

economic value and the heat used. The economic value

varies with the type of the energy source [46]. The

evaporation efficiency of a MD system is defined as the ratio

of the useful energy to the total energy [46,78] Useful energy

is the vapor diffusive energy i.e., energy required for the heat

of evaporation. The total energy is the sum of energies due to

conduction and vapor diffusion. The evaporation efficiency is given by [39,78]:

EE = =Q

diff

Qdiff

+ Qcond

=N H

v

mCp

T (52)

Where EE ( ) is the evaporation efficiency, Qdiff is the vapor diffusion energy and Qcond is the conductive energy.

The evaporation efficiency of the system increases with

the operating temperature. At high temperatures the heat

losses due to conduction become negligible, thereby

increasing the efficiency as shown in Fig. (13) [46]. Qtaishat

and co-wokers [38] observed that EE is low with low

operating temperatures. For such systems it is economical to

use low cost energy sources [68]. Burgoyne presented a

method where deaeration of the system increases the latent heat and reduces the heat losses to less than 10% [38].

Cabassud et al. stated that if the system is coupled with

solar energy and operated at higher temperature the energy to

be supplied can be maintained at low levels [89]. The

efficiency of the membrane can be maintained by cleaning

the membrane periodically to remove fouling layers.


6.6. Energy Requirements of Membrane Distillation

Process

Basic comparison has been made with currently available

desalination processes in terms of process and energy

requirements and is presented in Table 7.

Energy requirements and cost per production of unit

distillate for the membrane distillation process are compared

with other available desalination technologies using renew-able energy and summarized in Table 8 [90-93] below.

Since membrane distillation process does not require

high quality steam or hot feed solutions in the temperature

range 80-120oC as other distillation processes (MED and

MSF), it is sensible to utilize available low grade heat

sources. They could be direct solar energy harvested by solar

collectors, geothermal heat sources or any process waste heat

sources in the temperature range 40-80oC. MD coupled with

geothermal water sources has less energy requirements

compared to most other thermal desalination processes such

as MSF, MED and solar stills [97]. Membrane distillation

process has the potential to compete with reverse osmosis

provided a suitable membrane is employed in lower tempe-rature range. With suitable membrane of sufficient permea-

bility, vacuum membrane distillation process can compete

with RO at lower operating temperatures at 25oC [89].

Another advantage of the membrane distillation is that its

performance is not inhibited by high solute concentration

which is a big drawback for the reverse osmosis process

[10].

Fig. (13). Heat efficiency and loss vs. feed temperature. Permeate temperature at 20oC, Re = 4000 and distance at 0.35cm [45].

Table 7. Comparison of Existing Desalination Technologies

Parameter/Characteristic of the Process Membrane Distillation Reverse Osmosis Thermal Desalination

Type of process Membrane + Thermal Membrane Thermal

Membrane pore size 0.2-0.6 m 0.1-3.5 nm -

Feed temperature 60-85oC - 60-120oC

Cold water stream 20-25oC - May be required

Energy Thermal and mechanical Only mechanical Thermal and mechanical

Form of energy Low grade/waste heat from

renewable sources and some

mechanical energy for pumping

Requires prime quality

mechanical energy derived

from fossil fuels or

renewable sources

Steam, Low grade heat or waste heat

and some mechanical energy for

pumping

Fouling Less severe compared to RO due

to larger pore sizes

Easy fouling due to small

size pores [105]

Fouling at high temperatures

��

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��'

��(

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��

'��

*��

,��

!��

�


Some efforts to combine the membrane distillation with

renewable energy sources and long term performance

evaluations are summarized below in Table 9 [94-100].

7. EFFECT OF OPERATING CONDITIONS ON

PERMEATE FLUX

7.1. Effect of Feed Velocity

Feed velocity is an important operating condition in a

MD process as it affects the temperature difference across

the membrane and reduces the polarization at the membrane

surface. As shown in Fig. (14), at a given temperature, the

permeate flux increases with the feed velocity. This is

attributed to increases in the heat and mass transfer

coefficients at a higher Reynolds number [11, 46, 78, 101].

As a result, temperature and concentration polarization reduces at the boundary sides of the bulk streams. The flux

rise may also be due to a shorter retention time of the stream

in the membrane module leading to a higher transmembrane

temperature difference [47]. The permeate flux increases by

increasing the flow rate in the cold side of the membrane.

The effect of the flow rate in the permeate side is smaller as

compared to the effect in the feed side [46,76]

7.2. Effect of Feed Temperature

Temperature is one operating variable that affects not

only the MD process but also energy efficiency (EE)

significantly. Permeate flux increases with the temperature of the feed due to the exponential increase in the vapor

pressure with temperature [50, 68]. By increasing the inlet

temperature, the viscosity of the feed stream decreases [80]

and the temperature difference increases, generating more

heat removal from the feed solution. The effect of

temperatures in the range of 60-70oC on the permeate flux is

higher as compared to temperatures in the range of 40-50ºC)

[68].

The flux does not increase with temperature as rapidly as

the vapor pressure due the effect of the temperature polari-zation [88]. At a given flow rate, the effect of feed tempe-

rature on Reynolds number is very small [13]. Optimization

of feed flow rate and temperature parameters is important to

obtain effective water vapor flux.

The increase in the coolant temperature has reverse effect

on the permeate flux. High coolant temperature at constant

feed temperature reduces the temperature difference between

the hot and cold side and results in lower vapor pressure gradient as shown in Fig. (15) [73].

7.3. Effect of Concentration

Feed concentration has as a negative effect on the

performance and EE of the MD process. Increasing the salt

concentration in the feed, according to Raoult’s law causes a

decrease in the vapor pressure and in the pressure difference

(PF – PD), and therefore a decrease in the permeate flux as

shown in Fig. (16) [77, 87]. An extra concentration boundary

layer is developed adjacent to the feed membrane surface.

This layer together with the temperature boundary layer

decreases the driving force [9]. Also the flux decline is due

to the decrease in the heat and mass transfer boundary layer coefficients [45].

Higher feed concentrations result in higher mass transfer

resistance, lower energy efficiency and moderate heat trans-

fer resistance [78]. Higher feed concentrations also result in

the decline of the flux [71].

With high feed concentrations VMD process shows

better results compared to other processes [67]. Enhanced stirring rates reduced both the temperature and concentration

boundary layers and improved the DCMD performance [9].

7.4. Temperature Polarization

The temperature polarization coefficient (TPC) decreases

with the increase in the mean temperature (Tm). The energy

consumed for the vaporization of water in the feed increases

with the increase in Tm. As a result more heat flux takes

place through the liquid phase; hence resistance offered by

the heat transfer in liquid phase increases in relation to the

transmembrane mass transport. The consequence is the

increase in the temperature polarization and thus the flux

does not increase in proportion to the (dP/dT) [43,44,68]. Temperature polarization reduces with higher flow rates

because of a decrease in the thermal resistance to convection

in the boundary layer [50].

Table 8. Energy Requirements and Cost for Desalination Processes Using Renewable Energy

Process Energy Requirements

(kJ/kg)

Cost Per Unit Permeate

(US $/m3)

Year

Ref.

Membrane Distillation-Geothermal water (MD-GW) 111 15-18 2008 [90, 93]

Reverse Osmosis-Photovoltaic (RO-PV) 82 3.73 2002 [90, 93]

Multi Stage Flash Distillation (MSF) 338 2.02 1996 [90-92]

Multi Effect Distillation (MED) 240 2 1998 [90-92]

MESS (multi-effect solar still) 1500 12 2005 [90, 93]


Table 9. Summary of Applications of Renewable Energy in Membrane Distillation Process

ID Authors,

Year

Work Presented/Area of

study Experiments Results

Remarks and Outstanding

Contributions

1 Banat et al.

1994 [94]

The performance

characteristics of an AGMD

process for water

desalination were

investigated by conducting

two long-run experiments

The effect of hot side

temperature, cold side

flow rate and feed

concentration were

investigated in an 8-wk

and a 6-wk long run

experiments

The mass flux was steady over time.

The cold side flow rate had a

negligible effect on the permeate

flux and the flux exponentially

increased with increase in hot side

temperature.

The effects of temperature

gradient on mass transfer

was clarified by applying

the theory of non-

equilibrium

thermodynamics

2 Koschikows

ki et al.

2004 [95]

Theoretical simulations to

produce 0.2-20 m3/d.

experimental studies to

utilize solar collectors

Experimental setup with

solar-PV-power supply

system

Max. distillate production of 28 l/d

and min. of 16 l/d for m2 solar

collector area

Proved long term

operations without

maintenance using solar-

PV energy, energy

requirements were 117

kWh/m3 of distilled water

3 Gryta, 2002

[96]

Preparation of NaCl

supersaturated solutions by

direct contact membrane

distillation (DCMD)

integrated with salt

crystallization has been

studied

The feed temperature was

kept at

353 K or 358 K, and the

temperature of distillate

stream was changed from

293 K to 328 K

After 138 h of the process operation,

the permeate flux decrease

of about 40 % was observed

4 Banat et al.

2002 [97]

Integrating a membrane

distillation module with a

solar still

Indoor and outdoor

experiments conducted to

measure the sensitivity of

the permeate flux to the

brine temperature, flow

rate, salt concentration

and solar irradiation

Flux of water from the solar still was

no more than 20% of the total flux

The brine temperature

significantly affected the

flux of both the solar still

and the membrane module,

while the effect of salt

concentration was

marginal

5 Banat et al.

2007 [98]

Small-scale stand-alone

solar-driven (solar thermal

and PV energy) membrane

desalination unit was

designed and manufactured

Spiral-wound air-gap

membrane distillation

module with internal heat

recovery function was

used. Energy from

flat-plate solar thermal

collectors were used

120 L/d with an approximate

distillate conductivity of 5 μS/cm.

The thermal energy required by the

process was in the range of 200–300

kWh/m3

Proved feasibility of using

solar-PV energy for

membrane distillation

6 Banat et al.

2008 [99]

Exergy analysis method was

employed to evaluate the

exergy efficiency of the

“compact” and “large”

solardriven MD desalination

units

The exergy efficiency of

the compact and large

units with reference to the

exergy collected by the

solar collector was about

0.3% and 0.5%

The exergy efficiency of the flat

plate solar collectors in both units

varied diurnally and the maxima was

6.5% and 3% for the compact and

large units, respectively.

The highest exergy

destruction was found to

occur within the membrane

distillation module

7 Walton

et al., 2004

[100]

AGMD system was tested

using low-grade thermal

energy (13-75 ºC) supplied

by a salt-gradient solar

pond

Measured the flux per unit

area of membrane surface

and separation of ions

over a range of feedwater

salinities and temperature

as well as an assessment

of membrane fouling

Hot side temperatures as low as 13

ºC were tested. Flux per unit

temperature drop at very low

temperatures was only reduced by

50% compared to flux at higher

temperatures.

Operation of MD at such

low temperatures may

open up thermal energy

resources that have not

previously been considered

for desalination

8 Bouguecha

et al. 2005

[93]

Experimental studies were

conducted on the unit

combining membrane

distillation with geothermal

feed source

AGMD process has been

tested

Energy requirements for this process

were found to be 111 kJ/kg and the

cost for production is $13-18/m3

This configuration

compared to RO-PV in

terms of production cost


Fig. (14). Effect of feed side flow rate on flux for VF = VD, TF in = 353K and TD in = 293K. Average brine concentration 250-260g NaCl/dm3.

Solid and dotted lines represent various correlations [71].

Fig. (15). Effect of feed temperature on the water permeate flux. Average brine concentration 250-260g NaCl/dm3. Distillate inlet

temperature TD in = 293K. vF = 0.3m/sec and vD = 0.1m/sec. Solid dotted lines represent various correlations [71].

Temperature polarization coefficient (TPC) increases

with concentration. It is because the heat transfer coefficient decreases slightly with concentration and temperature

gradient in the liquid boundary layer decreases as lower

mass fluxes result in lower heat fluxes through the liquid

phases Fig. (17a) and Fig. (17b) [50].

As the Reynolds number increases, the TPC increases.

Cross flow has larger influence on the TPC than parallel

flow because high water vapor flux exists [13].

TPC increases with increase in recirculation rate. The

flux (mass or heat) does not increase with high TPC as the membrane distillation coefficient is reduced [44].

7.5. Concentration Polarization

In many cases the effect of feed concentration is smaller

as compared to the effect of the temperature polarization.

When highly concentrated salts are used the effects of

concentration polarization should be considered [85].

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The concentration polarization decreases the partial

vapor pressure of water at the feed-membrane interface;

because of this, the flux is reduced as the driving force for

evaporation is reduced.

7.6. Effect of Vapor Pressure Difference

The flux increases linearly with the increase in the vapor

pressure difference. The flux can be related to vapor pressure

difference by [13, 46, 50]:

N = p (53)

where is the membrane characteristic coefficient. depends on the operating conditions, permeate properties,

membrane characteristics, and module configurations. The equation is more accurate at low temperatures than at high temperatures as shown in Fig. (18) [46].

7.7. Effect of Heat Transfer Coefficient

The heat transfer coefficients (h) increase with the increase in mean temperature, thus increasing the molar flux. The effect of temperature on the feed side heat transfer coefficient is high compared to the permeate side heat trans-fer coefficient. This is because the variations of temperature on the feed side is more significant than on the permeate side [68]. The h values increase with the recirculation rates [43].

Fig. (16). The effect of feed concentration on permeate flux (points). Lines are model predictions. TF in = 353K and TD in = 293K [71]

(a) (b)

Fig. (17). (a) Temperature polarization coefficients for various feed temperatures in laminar flow for three types of membranes [49];(b)

Temperature polarization coefficients for various feed temperatures in turbulent flow for three types of membranes [50].

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Fig. (18). Flux as function of vapor pressure difference for per-

meate temperature 200C, Volumetric flow rate 5.5 L/min and length

0.35cm [46].

The presence of spacers increase the film transfer coeffi-cients, thus increasing the flux [88]. For highly polarized systems, the increase in film transfer coefficients can enhance the flux [87]. Figures (19a) and (19b) show the relation between the temperature and the boundary layer’s heat transfer coefficient and conductive heat transfer coefficients respectively.

7.8. Effect of Membrane Mass Transfer Coefficient

With increase in the feed bulk temperature, the mass transfer coefficient (C) also increases due to increase in the water vapor diffusion coefficient through the stagnant air as shown in Fig. (20) [68]. With the increase in the recirculation

rates, the C value decreases because of the increase in the air pressure in the pores [43]. The reduction in C when spacers are used is a consequence of change in area of membrane for evaporation and condensation as some part of area is obscured at the point of contact between spacer and mem-brane [44]. Phattaranawik et al. observed that mass transfer coefficient does not change with flow rate and temperature within the experimental range because the average temperature change was less than 5% and also mass transfer occurred in transition region [50].

Fig. (20). Theoretical and experimental values of membrane mass

transfer coefficient vs. bulk feed temperature [68].

8. APPLICATIONS OF MEMBRANE DISTILLATION

Membrane distillation was first designed to de-mineralize brackish water at a desirable low energy consumption to compete with conventional evaporation processes. Besides water purification, it has now been widely used in many chemical engineering processes, including solution con-centration, and membrane crystallization, with either hydrophobic or hydrophilic membrane.

8.1. Water Purification/Treatment

Over 90% of membrane distillation patents relates to water desalination. In principle, the process rejection of non-

Fig. (19). (a) Effect of average bulk temperature of feed and permeate streams on boundary layer’s heat transfer coefficient with distilled

water as feed and permeate liquid and flow rate vF = vP = 2 L/min [66] (b) Effect of average bulk temperature of feed and permeate streams

on conductive heat transfer coefficient of membrane, hm with distilled water as feed and permeate liquid and flow rate vF = vP = 2 L/min [68]

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volatile dissolved compounds is 100%. Therefore, it is possible by such process to produce high quality pure water, fully demineralized, or drinking water from seawater or underground brackish water. Membrane distillation is also used to treat wastewater containing heavy metal ions or radioactive elements or other wastes, with the purpose of reducing the waste volume.

Fane and colleagues [102] constructed a simple mem-brane bioreactor for treating fermented water. The inflow wastewater was first degraded by biological agents in the bioreactor; further distillation was conducted to permeate across a membrane. Membrane configurations such as DCMD, AGMD, AGMD, and VMD were carried out in the membrane bioreactor. Compared to other membrane processes, such as MF, UF and NF, MD was found more effective in removing metabolic organics from permeate, requiring less energy consumption.

8.2. Solution Concentration

Solution concentration is quite similar to water purifi-cation process presented in section 8.1, with the only difference that the concentrated outflow solution from the feed stream side is now the desired product. Since it can be operated at room temperature, it is extremely useful in cases where the solution is heat sensible, such as food and biological products.

Ramakrishnan et al. [103] used a DCMD process to concentrate hydroxitic acid with Garcinia extraction. Using a polypropylene flat plate membrane at 25°C and ambient pressure, a final 5-6 times concentrated solution of HCA was obtained without any energy loss. Shaw [104] described methods for using membrane distillation to concentrate biological solutions such as proteins or DNAs or RNAs and other macromolecules. Two-layer membranes were used to enhance the vapor flux. The first layer was water-vapor permeate permitted, while the other layer was liquid-water wet permitted.

It has been suggested the coupling of MD for processes in which a product/substrate aqueous solution needs to be concentrated while another product/substrate aqueous solution needs to be diluted with water. Drackett et al. [105] used MD in chlorine-alkali industries. In this process, nearly saturated NaCl solution is fed to an electrolyzer to produce chlorine gas at the anode while water is reduced to hydrogen and hydroxyl ions at the cathode. In order to shorten the steps of re-concentrating the spent NaCl solution to the original feed concentration, MD is introduced to integrate NaCl solution stream from the electrolyzer with the NaOH solution stream from the catholyte. By this means, NaCl brine is re-concentrated and can be directly re-fed to the electrolyzer.

8.3. Membrane Crystallization

Membrane crystallization can be applied in the prepa-ration of single crystals. In theory, the crystallization process is simple. A target compound in pure form is dissolved in solvent. The chemical environment of the dissolved target material is then altered so that the target is less soluble and reverts to the solid phase in crystalline form. In practice,

however, forming a high quality single crystal is rather difficult, requiring trial and error steps and much patience of the researcher. The main problem deals with the environment alteration, such as adding a crystal seed by changing either the temperature or pressure to avoid the formation of multi-crystals. Ake’s research group [106, 107] proposed the use of MD to obtaining a protein single crystal in which crystal growth housing is separated by a permeable membrane from an adjacent well having a much larger volume. The well is configured to contain a fluid having identity and concen-tration similar to the solvent and crystallizing agent emp-loyed in crystal growth. By slightly changing the chemical environment in the well, the chemical environment in the crystal growth housing will change even more slightly due to the mass and heat transfer across the membrane. MD pro-mises a more favorable procedure to obtain a single crystal.

9. CURRENT STATUS & PERSPECTIVES

Compared to conventional evaporation and other membrane techniques, membrane distillation is claimed as a capital and energy saving technology with many advantages. It is supposed to be the most promising and emerging technique for purification of saline water. The potential to apply this technique to other fields is also of particular interest and under development [1, 2]. Although membrane distillation has been investigated in research labs for more than 40 years and numerous research articles and patents are published worldwide, this technique is still not fully commercialized. Some pilot plants may have recently come up, but few full scale industrial applications can be found [108, 109]. This may be due to two main drawbacks in the process. The first one is related to a low productivity. The current maximum permeate flux in the membrane distillation desalination process is about 80 kg m-2·h-1, with a low productivity of 1%-10% as compared to other membrane treatment process such as ultra-filtration or reverse osmosis [110, 111]. The second drawback is related to the presence of temperature polarization which results in additional energy loss. Current efforts and developments to overcome these two problems include: (1) using new membrane materials; (2) developing novel structures for membrane modules; and (3) combining MD with other separation or reaction processes.

To gain special application benefits and to save energy, MD has been combined with some other processes. Colling et al. [112] coupled membrane distillation with fractional distillation to separate olefins from paraffins. The membrane has been placed on the top of the distillation column (re-flux stage), around the column (distillation stage), or at the bottom of the column. It has been stated that the preferred location of the membrane may have an effect on the driving force of the mixture [113,114]. Huff et al. [115] and Miller et al. [116] combined membrane distillation with simulated moving bed and fractional crystallization to purify mixture of xylene. Tian et al. used solar energy to produce fresh water through membrane distillation with higher energy efficiency [117].

CURRENT & FUTURE DEVELOPMENTS

The membrane distillation process, a thermally driven process, although still under experimental and pilot-plant


stage will be a viable alternative to membrane and thermal processes such as reverse osmosis and conventional distillation. It is expected to be a complete desalination pro-cess in the near future. Long term performance is crucially required for the industrial application of membrane distil-lation. The major reason that restricts the MD process to be implemented in the industry is its flux decay due to deposit formation. The deposit formation can be controlled by an integrated use of the membrane process.

Currently, membranes used in microfiltration are being used for membrane distillation. The characteristics and properties of the membranes play a vital role on the molar permeate flux. Studies on careful optimization of membrane materials and membrane modules to reduce the polarization effects are gaining importance. Only few authors have considered the possibility of manufacturing new membranes and membrane modules specifically for MD process. Composite membrane preparations, using surface coating, fabrication of dual layer membranes, porous hydrophobic/ hydrophilic membranes are some recent improvements in the membrane material developments. The membrane distil-lation unit is said to be performing best if high flux is obtained at minimal energy consumption. Flux is observed to decay with time because of pore wetting, fouling and scaling problems. Lack of sufficient data on membrane fouling and scaling require more experimental studies to be carried out. The membranes should be chemically resistant to cleaning agents.

Kinetic studies and modeling of process under variable characteristics are important and these are to be considered while optimizing the process. Optimizing the process by increasing the flux, decreasing energy consumption and providing long term performance will make the process a viable alternative to many applications.

ACKNOWLEDGEMENTS

We appreciate the financial supports from ConocoPhillips R&D, and the postdoctoral fellowship for Dr. Gude provided by New Mexico State University (College of Engineering, Chemical Engineering Department, and Civil & Environmental Engineering Department).

CONFLICT OF INTEREST

The authors declare no conflict of interest.

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membrane distillation for desalination and other separations

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