An Introduction to Membrane Distillation

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An Introduction to Membrane
Distillation
Christos Charisiadis 2014
Membrane distillation: A comprehensive review
Desalination 287 (2012) 2–18
Advances in Membrane Distillation for Water Desalination and Purification Applications
Water 2013, 5, 94-196; doi:10.3390/w5010094

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Contents
1. Introduction 4
1.1 Using MD has many attractive features 4
1.2 MD is also attended by some drawbacks 5
1.3 Modelling Aspects of Membrane Distillation 5
2. Membrane configuration 6
2.1. Direct Contact Membrane Distillation (DCMD) 6
2.2. Air Gap Membrane Distillation (AGMD) 7
2.3. Sweeping Gas Membrane Distillation (SGMD) 7
2.4. Vacuum Membrane Distillation (VMD) 8
3. Configurations of MD Modules 9
4. Membranes for MD Applications 10
4.1 Hollow fiber 10
4.2 Plate and frame 11
4.3. Tubular membrane 11
4.4. Spiral wound membrane 11
5. Membrane Material 12
6. Membrane characteristics 12
6.1 Membrane thickness 13
6.2 Liquid entry pressure (wetting pressure) 13
6.3 Membrane porosity and tortuosity 15
6.4 Mean pore size and pore size distribution 15
6.5 Thermal conductivity 16
6.6 Surface Roughness 17
7. Advances on MD Processes and Modules for Water Purification 18
7.1. MD Stand-Alone Systems 18

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7.2. State of the Art MD Research and Systems 19
8. Mechanisms 20
8.1 Mass transfer 21
8.1.1 Direct Contact Membrane Distillation (DCMD) 21
8.1.2. Air Gap Membrane Distillation (AGMD) 24
8.1.3. Vacuum Membrane Distillation (VMD) 26
8.1.4. Sweeping Gas Membrane Distillation (SGMD) 27
8.2. Heat transfer 28
9. Thermal efficiency and energy consumption 33
10. Temperature polarization and concentration polarization 34
11. Fouling 35
12. Operating parameters 37
12.1 Parameters to Reducing Temperature Polarization 37
12.2 Feed temperature 38
12.3 The concentration and solution feature 39
12.4 Recirculation rate 40
12.5 The air gap 40
12.6 Membrane type 41
13. Future Developments and Conclusions 41

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1. Introduction
Membrane Distillation (MD) is a thermally-driven separation process, in which only
vapour molecules transfer through a microporous hydrophobic membrane. The
driving force in the MD process is the vapour pressure difference induced by the
temperature difference across the hydrophobic membrane. This process has various
applications, such as desalination, wastewater treatment and in the food industry.
1.1 Using MD has many attractive features
1. Low operating temperatures in comparison to those encountered in conventional
process; the solution (mainly water) is not necessarily heated up to the boiling point.
2. The hydrostatic pressure encountered in MD is lower than that used in pressure-
driven membrane processes like reverse osmosis (RO). Therefore, MD is expected to
be a cost effective process, which requires less demanding of membrane
characteristics too. In this respect, less expensive material can be involved in it such
as plastic, for example, thus alleviating corrosion problems.
3. According to the principle of vapour–liquid equilibrium, the MD process has a high
rejection factor. As a matter of fact, theoretically, complete separation takes place.
In addition, the membrane pore size required for MD is relatively larger than those
for other membrane separation processes, such as RO. The MD process, therefore,
suffers less from fouling.
4. The MD system has the feasibility to be combined with other separation processes
to create an integrated separation system, such as ultrafiltration (UF) or with a RO
unit.
5. MD has the ability to utilize alternative energy sources, such as solar energy.
The MD process is competitive for desalination of brackish water and sea water. It is
also an effective process for removing organic and heavy metals from aqueous
solution, from waste water. MD has also been used to treat radioactive waste,
where the product could be safely discharged to the environment.

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1.2 MD is also attended by some drawbacks
1. Low permeate flux (compared to other separation processes, like RO)
2. High susceptibility permeate flux to the concentration and temperature of the
feed conditions due to the concentration and temperature polarization
phenomenon.
3. The trapped air within the membrane introduces a further mass transfer
resistance, which also limits the MD permeate flux.
4. The heat lost by conduction is quite large.
1.3 Modelling Aspects of Membrane Distillation
Mass transfer in MD is accompanied by heat transfer, so MD Modelling is more
complex than that for heat exchangers. The parameters or data that should be
considered during Modelling include:
• Membrane characteristics, such as membrane thickness, pore size, tortuosity and
porosity, which can be aquired by gas permeation experiment, scanning electron
microscopy and image analysis;
• Thermal conductivity of the membrane is measurable in some cases or can be
calculated;
• Convective heat transfer coefficient of the feed and/or permeate streams, which
can be calculated by semi-emperical equations based on Nusselt numbers and by
including factors such as the structure of the spacer or module, flow velocities,
properties of feed and permeate, the operation temperature, etc. and
• An important assumption adopted in Modelling MD is that the kinetic effects at the
vapor-liquid interface are negligible. According to this assumption, vapor-liquid
equilibrium equations can be applied to determine the partial vapor pressures of
each component at each side of the membrane.
Mass transfer in MD is controlled by three basic mechanisms,
1. Knudsen diffusion
2. Poiseuille flow (viscous flow)
3. Molecular diffusion.
This gives rise to several types of resistance to mass transfer resulting from transfer
of momentum to the supported membrane (viscous), collision of molecules with
other molecules (molecular resistance) or with the membrane itself-(Knudsen

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resistance (see Fig. 1). In this context, the dusty gas model is used to describe the
mass transfer resistances in the MD system. It is worth mentioning that the mass
transfer boundary layer resistance is generally negligible.
Similarly, the surface resistance is insignificant, because the surface area of the MD
is small compared to the pore area. On the other hand, the thermal boundary layer
is considered to be the factor limiting mass transfer.
2. Membrane configuration
In this section, different MD configurations that have been utilized to separate
aqueous feed solution using a microporous hydrophobic membrane will be
addressed.
2.1. Direct Contact Membrane Distillation (DCMD)
In this configuration (Fig. 2), the hot solution (feed) is in direct contact with the hot
membrane side surface. Therefore, evaporation takes place at the feed-membrane
surface. The vapour is moved by the pressure difference across the membrane to the
permeate side and condenses inside the membrane module. Because of the
hydrophobic characteristic, the feed cannot penetrate the membrane (only the gas
phase exists inside the membrane pores). DCMD is the simplest MD configuration,
and is widely employed in desalination processes and concentration of aqueous
solutions in food industries, or acids manufacturing. The main disadvantage for
DCMD in commercial applications is its low energy efficiency. Although polymeric

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membranes generally have low thermal conductivity, the driving force (temperature
difference between the feed and permeate sides) for mass transfer will also lead to
significant conductive heat transfer through the membrane due to the small
membrane thickness, so only part of the supplied heat energy is used for production.
Of the four configurations, DCMD has the highest heat conduction loss because of
the higher heat transfer coefficient on the permeate side for this configuration,
which results in relatively low thermal efficiency.
2.2. Air Gap Membrane Distillation (AGMD)
The schematic of the Air Gap Membrane Distillation (AGMD) is shown in Fig. 3. The
feed solution is in direct contact with the hot side of the membrane surface only.
Stagnant air is introduced between the membrane and the condensation surface.
The vapour crosses the air gap to condense over the cold surface inside the
membrane cell. The benefit of this design is the reduced heat lost by conduction.
However, additional resistance to mass transfer is created, which is considered a
disadvantage. This configuration is suitable for desalination and removing volatile
compounds from aqueous solutions.
In AGMD, the air gap is usually the controlling factor for the mass and heat transfers
because of its greater thermal and mass transfer resistances. In comparison with the
thickness (40–250 μm) and conductivity of the membrane, the air gap is much
thicker (general 2000–10,000 μm) and has lower thermal conductivity. Therefore,
more heat energy in AGMD will be used for water evaporation than in DCMD.
Additionally, if a low temperature feed is used as the cooling stream in this
configuration, the latent heat can be recovered through the condensation of the
vapor on the cooling plate. However, AGMD typically has a low flux, due to the low
temperature difference across the membrane and therefore larger surface areas are
required.
2.3. Sweeping Gas Membrane Distillation (SGMD)
In Sweeping Gas Membrane Distillation (SGMD), as the schematic diagram in Fig. 4
shows, inert gas is used to sweep the vapour at the permeate membrane side to
condense outside the membrane module. There is a gas barrier, like in AGMD, to

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reduce the heat loss, but this is not stationary, which enhances the mass transfer
coefficient. This configuration is useful for removing volatile compounds from
aqueous solution. The main disadvantage of this configuration is that a small volume
of permeate diffuses in a large sweep gas volume, requiring a large condenser.
In SGMD, the vapor is stripped from the hot feed by a gas stream, and then
condensed in an external condenser. It has higher mass transfer rates than AGMD,
due to the greater driving force originating from the reduced vapor pressure on the
permeate side of the membrane, and has less heat loss through the membrane than
DCMD. However, an external condenser and an air blower or compressed air are
needed to maintain operation of this configuration, which will cause an increase in
investment, energy use and running costs.
It is worthwhile stating that AGMD and SGMD can be combined in a process called
thermostatic sweeping gas membrane distillation (TSGMD). The inert gas in this case
is passed through the gap between the membrane and the condensation surface.
Part
of
vapour is condensed over the condensation surface (AGMD) and the remainder is
condensed outside the membrane cell by external condenser (SGMD).
2.4. Vacuum Membrane Distillation (VMD)
The schematic diagram of this module is shown in Fig. 5. In VMD configuration, a
pump is used to create a vacuum in the permeate membrane side. Condensation
takes place outside the membrane module. The heat lost by conduction is negligible,
which is considered a great advantage. This type of MD is used to separate aqueous
volatile solutions.

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In VMD, the vapor permeate is removed continuously from the vacuum chamber to
form a vapor pressure difference across the membrane. Theoretically, this
configuration can provide the greatest driving force at the same feed temperature,
because the vapor pressure on the cold side can be reduced to almost zero. An
external condenser is required as for AGMD, if the liquid permeate is the product.
Of the four configurations, DCMD is the most popular for MD laboratory research,
with more than half of the published references for membrane distillation based on
DCMD. However, AGMD is more popular in commercial applications, because of its
high energy efficiency and capability for latent heat recovery
3. Configurations of MD Modules
There are two major MD module configurations, which are the tubular module and
the plate and frame module. Both of these modules have been used in pilot plant
trials.
Figure (a) shows a schematic diagram of a hollow fiber tubular module, in which
hollow fiber membranes are glued into a housing. This configuration can have a very
high packing density (3000 m2
/m3
). The feed is introduced into the shell side or into
lumen side of the hollow fibers, and cooling fluid, sweeping gas, or negative pressure
can be applied on the other side to form VMD, SGMD, or DCMD. Because of its large
active area combined with a small footprint, hollow fiber modules have great
potential in commercial applications. Although broken hollow fibers cannot be
replaced, they can be detected by the liquid decay test (LDT) and pinned to remove
broken fibers from service. Good flow distribution on the shell side can be difficult to
achieve, with subsequent high degrees of temperature polarization. Cross-flow
modules have been developed to reduce this effect for hollow fiber modules
Figure (b) shows the structure of the plate and frame module. This module is suitable
for flat sheet membranes and can be used for DCMD, AGMD, VMD, and SGMD. In
this configuration, the packing density is about 100–400 m2
/m3
. Although this

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configuration has a relatively smaller effective area for the same volume when
compared to the tubular modules, it is easy to construct and multiple layers of flat
sheet MD membranes can be used to increase the effective area. As shown in Figure
(b), it is easy to change damaged membranes from this configuration. Thus, this
module is widely employed in laboratory experiments for testing the influence of
membrane properties and process parameters on the flux or energy efficiency of
membrane distillation. Also the flow dynamics can be improved by the use of spacers
that increase turbulence and reduce temperature polarization.
To meet the requirement of commercial applications, other configurations with large
specific areas were also developed, i.e., spiral-wound modules mainly employed for
air/permeate gap membrane distillation have a much more compact structure than
the conventional plate and frame AGMD.
4. Membranes for MD Applications
There are two common types of membrane configurations shown in Figure 3:
• Hollow fiber membrane mainly prepared from PP, PVDF, and PVDF-PTFE composite
material; and
• Flat sheet membrane mainly prepared from PP, PTFE, and PVDF.
4.1 Hollow fiber
The hollow fiber module, which has been used in MD, has thousands of hollow fibers
bundled and sealed inside a shell tube. The feed solution flows through the hollow
fiber and the permeate is collected on the outside of the membrane fiber (inside-
outside), or the feed solution flows from outside the hollow fibers and the permeate
is collected inside the hollow fiber (outside-inside). The main advantages of the
hollow fiber module are very high packing density and low energy consumption. On
the other hand, it has high tendency to fouling and is difficult to clean and maintain.
It is worth mentioning that, if feed solution penetrates the membrane pores in shell
and tube modules, the whole module should be changed.

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Compared with flat sheet membranes, hollow fiber membranes have relatively large
specific surface areas, but the main impediment of the hollow fiber module is its
typically low flux (generally 1–4 L m−2
h−1
at 40–60 °C). The low flux is related to its
poor flow dynamics and the resultant high degree of temperature polarization.
However, high-flux hollow fiber membranes with different features suitable for
membrane distillation have been developed recently, such as dual-layer hydrophilic-
hydrophobic fibers with a very thin effective hydrophobic PVDF layer (50 μm), and
hollow fiber membranes with a sponge-like structure and thin walls, which have flux
of about 50–70 kg m−2
h−1
at about 80–90 °C. This flux is as high as that from flat
sheet membrane.
4.2 Plate and frame
The membrane and the spacers are layered together between two plates (e.g. flat
sheet). The flat sheet membrane configuration is widely used on laboratory scale,
because it is easy to clean and replace. However, the packing density, which is the
ratio of membrane area to the packing volume, is low and a membrane support is
required.The flat sheet membrane is used widely in MD applications, such as
desalination and water treatment.
The reported flux from flat sheet membranes is typically 20–30 L m−2
h−1
at inlet
temperatures of hot 60 °C and cold 20 °C. In general, the polymeric membrane
shown in Figure 3b is composed of a thin active layer and a porous support layer.
This structure is able to provide sufficient mechanical strength for the membrane to
enable the active layer to be manufactured as thin as possible, which reduces the
mass transfer resistance.
4.3. Tubular membrane
In this sort of modules, the membrane is tube-shaped and inserted between two
cylindrical chambers (hot and cold fluid chambers). In the commercial field, the
tubular module is more attractive, because it has low tendency to fouling, easy to
clean and has a high effective area. However, the packing density of this module is
low and it has a high operating cost. Tubular membranes are also utilized in MD.
Tubular ceramic membranes were employed in three MD configurations DCMD,
AGMD and VMD to treat NaCl aqueous solution, where salt rejection was more than
99%.
4.4. Spiral wound membrane
In this type, flat sheet membrane and spacers are enveloped and rolled around a
perforated central collection tube. The feed moves across the membrane surface in
an axial direction, while the permeate flows radially to the centre and exits through
the collection tube. The spiral wound membrane has good packing density, average

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tendency to fouling and acceptable energy consumption. It is worth stating that
there are two possibilities for flow in a microfiltration system; cross flow and dead-
end flow. For cross flow, which is used in MD, the feed solution is pumped
tangentially to the membrane. The permeate passes through the membrane, while
the feed is re-circulated. However, all the feed passes through the membrane in the
dead-end type.
5. Membrane Material
The most common materials used for MD membranes are polytetrafluoroethylene
(PTFE), polypropylene (PP) and polyvinylidenefluoride (PVDF). The porosity of the
membranes used is in the range of 0.60 to 0.95, the pore size is in the range of 0.2 to
1.0 μm, and the thickness is in the range of 0.04 to 0.25 mm. The surface energies
and thermal conductivities of these materials are listed in Table 1.
Of these materials, PTFE has the highest hydrophobicity (largest contact angle with
water), good chemical and thermal stability and oxidation resistance, but it has the
highest conductivity which will cause greater heat transfer through PTFE
membranes. PVDF has good hydrophobicity, thermal resistance and mechanical
strength and can be easily prepared into membranes with versatile pore structures
by different methods. PP also exhibits good thermal and chemical resistance.
Recently, new membrane materials, such as carbon nanotubes, fluorinated
copolymer materials and surface modified PES, have been developed to make MD
membranes with good mechanical strength and high hydrophobicity and porosity.
6. Membrane characteristics
In membrane distillation, membranes on the basis of their selective properties are
not involved in the mass transport phenomena, but are involved in heat transport
from the hot side to the cold side. Therefore, compounds transferred across the
membrane in gas phase are driven by vapour pressure differences based on vapour-
liquid equilibrium, and the macro-porous polymeric or inorganic membrane
employed between the permeate and feed sides acts as a physical barrier providing
the interfaces where heat and mass are simultaneously exchanged.
A suitable membrane needs to exhibit certain characteristics in order to be viable in
MD. Although, the membrane should be porous, it should not be wetted by the

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process liquids under the pressure applied within the membrane module.
Furthermore, no capillary condensation should occur within the pores, while the
membrane itself should not affect the vapour/liquid equilibrium of the system being
desalinated. Finally, at least one side of the membrane should be in direct contact
with the process liquids while only vapour should be transported across the pores of
the membrane.
The properties and structure of the surface of the membrane are highly important,
and the requirements may vary depending on application and the type of MD
configuration in which the membrane is being used: VMD, AGMD, SGMD or DCMD.
While SGMD and VMD are the most energy intensive MD techniques, they are
generally preferred to separate two mixed liquids having different boiling points to
avoid further treatments linked to the contamination of another carrier liquid on the
permeate side. On the other hand for water purification, desalination and
dewatering AGMD and DCMD are generally used since only water is evaporated from
the bulk feed.
Thus, the properties of membranes suitable for membrane distillation should
include:
6.1 Membrane thickness
An adequate thickness, based on a compromise between increased membrane
permeability (tend to increase flux) and decreased thermal resistance (tend to
reduce heat efficiency or interface temperature difference) as the membrane
becomes thinner;
The membrane thickness is a significant characteristic in the MD system. There is an
inversely proportional relationship between the membrane thickness and the
permeate flux. The permeate flux is reduced as the membrane becomes thicker,
because the mass transfer resistance increases, while heat loss is also reduced as the
membrane thickness increases. The optimum membrane thickness lies between 30–
60 μm. It is worth noting that the effect of membrane thickness in AGMD can be
neglected, because the stagnant air gap represents the predominant resistance to
mass transfer.
6.2 Liquid entry pressure (wetting pressure)
Reasonably large pore size and narrow pore size distribution, limited by the
minimum Liquid Entry Pressure (LEP) of the membrane. In MD, the hydrostatic
pressure must be lower than LEP to avoid membrane wetting. This can be quantified
by the Laplace (Cantor) Equation as following Equation (1):
(1)

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where Pf and Pp are the hydraulic pressure on the feed and permeate side, B is a
geometric pore coefficient (equal to 1 for cylindrical pores), γl is liquid surface
tension, θ contact angle and rmax is the maximum pore size. The contact angle of a
water droplet on a Teflon surface varies from 108° to 115°; 107° for PVDF and 120°
for PP. It is worthwhile indicating that a flat ceramic membrane had a contact angle
varying from 177° to 179°. Moreover, ceramic zirconia and titania hydrophobic
membranes were prepared with a 160° contact angle. All these ceramic membranes
are used for desalination purposes (see Table 1). With regard to surface tension, the
impact of salt concentration (NaCl) on the water surface tension is:
(2)
γl stands for pure water surface tension at 25 °C, which is 72 mN/m
As a result, membranes that have a high contact angle (high hydrophobicity), small
pore size, low surface energy and high surface tension for the feed solution possess a
high LEP value. Typical values for surface energy for some polymeric materials are
reported in Table 2 below. The maximum pore size to prevent wetting should be
between 0,1–0,6 μm. Moreover, the possibility of liquid penetration in VMD is higher
than other MD configurations, so a small pore size is recommended.

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6.3 Membrane porosity and tortuosity
High porosity. High porosity increases both the thermal resistance and the
permeability of MD membranes, so both the heat efficiency and flux are increased.
However, high porosity membranes have low mechanical strength and tend to crack
or compress under mild pressure, which results in the loss of membrane
performance.
Membrane porosity refers to the void volume fraction of the membrane (defined as
the volume of the pores divided by the total volume of the membrane). Higher
porosity membranes have a larger evaporation surface area. Two types of liquid are
used to estimate membrane porosity. The first penetrates the membrane pores (e.g.
isopropyl alcohol, IPA), while the other, like water, does not. In general, a membrane
with high porosity has higher permeate flux and lower conductive heat loss. The
porosity (ε) can be determined by the Smolder–Franken equation.
(3)
where ρm and ρpol are the densities of membrane and polymer material, respectively.
Membrane porosity in the MD system varies from 30 to 85%.
Tortuosity (τ) is the deviation of the pore structure from the cylindrical shape. As a
result, the higher the tortuosity value, the lower the permeate flux. The most
successful correlation is:
(4)
6.4 Mean pore size and pore size distribution
Low surface energy, equivalent to high hydrophobicity, based on Equation (1).
Material with higher hydrophobicity can be made into membranes with larger pore
sizes, or membranes made from more hydrophobic material will be applicable under
higher pressures for a given pore size;
Membranes with pore size between 100 nm to 1 μm are usually used in MD systems.
The permeate flux increases with increasing membrane pore size. The mechanism of
mass transfer can be determined, and the permeate flux calculated, based on the
membrane pore size and the mean free path through the membrane pores taken by
transferred molecules (water vapour). Generally, the mean pore size is used to
determine the vapour flux. A large pore size is required for high permeate flux, while
the pore size should be small to avoid liquid penetration. As a result, the optimum
pore size should be determined for each feed solution and operating condition.

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In fact, the membrane does not have a uniform pore size so more than mass transfer
mechanisms occur simultaneously (depending to the pore size). There are several
investigations examine the importance of pore size distribution in MD flux. Better
understanding of membrane morphology such as pore size, pore size distribution,
porosity, and thickness directs to have an accurate mass and heat transfer modeling.
Regarding to the MD membrane, two types of characteristics can be analyzed, the
structural characteristic and the actual separation parameters (permeation).
6.5 Thermal conductivity
Low thermal conductivity. High thermal conductivities increases sensible heat
transfer and reduce vapor flux due to reduced interface temperature difference;
The thermal conductivity of the membrane is calculated based on the thermal
conductivity of both polymer ks and gas kg (usually air). The thermal conductivity of
the polymer depends on temperature, the degree of crystallinity, and the shape of
the crystal. The thermal conductivities of most hydrophobic polymers are close to
each other. For example, the thermal conductivity of PVDF, PTFE and PP at 23 °C are
0.17–0.19, 0.25–0.27 and 0.11–0.16 Wm−1
K−1
respectively. The thermal conductivity
of PTFE can be estimated by
Ks = 4,86 x 10-4
x T + 0,253 (5)
The thermal conductivity of the MD membrane is usually taken a volume-average of
both conductivities ks and kg as follows:
km = (1-ε) x ks + ε x kg (6a)
However, it is suggested that thermal conductivity of an MD membrane is better
based on the volume average of both resistances (1/kg and 1/ks), i.e.,
km = [ e / kg + (1-ε) / ks]-1
(6b)
for, the thermal conductivity values for air and water vapour at 25 °C are of the same
order of magnitude. For instance, the thermal conductivity of air at 25 °C is
0.026Wm−1
K−1
and for water vapour, it is 0.020Wm−1
K−1
. As a result, the assumption
of one component gas present inside the pores is justified. The thermal conductivity
of water vapour and air at around 40 °C can be computed by:
kg = 1,5 x 10-3
x (7)
Some ways to reduce the heat loss by conduction through the membrane; using
membrane materials with low thermal conductivities, using a high porosity
membrane, using thicker membrane, and minimizing heat losses. It is also suggested
that the permeability can be enhanced by using a composite porous

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hydrophobic/hydrophilic membrane. In this case, the top layer is very thin
hydrophobic layer to stop liquid penetration, followed by thick hydrophilic layer.
Both layers reduce the heat losses through the membrane.
6.6 Surface Roughness
Surface roughness is also critical, because it will affect a number of properties
including surface fouling and the contact angle of water on the membrane surface. A
change of wetting behaviour will likely affect heat conduction across the top
membrane layer, therefore, clearly affecting performance of MD membranes.
Although wetting was shown to be facilitated by rough hydrophilic surfaces as more
points for spreading are offered to the liquid, this is not always true for
homogeneous hydrophobic materials and was shown to highly depend on the
composition of the surface and the shape of the roughness extrusions. As the
average roughness increases, the advancing angle of liquids on hydrophobic surfaces
tends to be increased due to the larger number of interactions between the nodules
and obstacles composing the surface of the membrane and the liquid. This tendency,
known in surface science as the lotus effect, is particularly enhanced for materials
exhibiting contact angles >150° with the wetting liquid. A convenient way to
measure roughness is typically given by the roughness factor κ, defined as:
κ = Αm / An (8)
where An and Am are respectively the area calculated as the projection of the object
on a plan normal to the main direction of the surface, and the surface area
measured by any experimental adsorption technique.
The measured surface roughness and area can be obtained by a number of
techniques, such as atomic force microscopy (AFM), gas adsorption (BET), diffuse X-
ray spectroscopy or laser light scattering depending on the size of the pores and the
accuracy sought. In the case of membrane surfaces, the difficulty resides in the
definition of what the true roughness is or, in other words, how deep one wants to
consider fluctuations from the surface as the surface or the inside of the pores. The
characterization of surface roughness is often ignored in MD as the process is
considered to be mostly unaffected by fouling. However, surface roughness as
shown does have other implications on the performance of membranes for MD and
should, therefore, be more thoroughly studied.
7. Advances on MD Processes and Modules for Water Purification
Even though membrane distillation was patented in the 1960s, it has not been
commercialised because of the success of competing technologies. However in just
the last few years, MD has emerged with numerous commercially oriented devices

18
and novel process integrations. This section focuses on the current process
arrangements and commercially available MD systems.
7.1. MD Stand-Alone Systems
A module to house a membrane and perform MD is not complicated but requires
more complexity in its connections as compared to pressurised membrane systems
(micro, ultra and nanofiltration as well as reverse osmosis). As shown in Figure 6, we
see the simplest form of DCMD configuration which will desalinate a saline water
feed to a very high quality permeate.
However, the simplest form suffers drawbacks which must be overcome to make MD
practically useful. The three key drawbacks under standard process configuration
are:
• Water recovery limit: The flux of the membrane draws a significant amount of
energy purely through the evaporation of the feed, which is deposited into the
permeate. The limiting amount of water permeated as a fraction of water fed, F,
(i.e., single pass recovery) is presented according to as Equation (9):
F = (1-t) x CP x (TF - TE)/ΔΗvap (9)
where TF and TE are the feed and exit temperatures, respectively (K or °C), CP is the
specific heat of water (4.18 kJ/kg/K), t is the proportion of conductive heat (balance
due to evaporative heat) loss through the membrane, and ΔHvap is the latent heat of
vaporisation (kJ/kg). For example, if the feed water is supplied at 80 °C, no more
than 7.7 wt % of this desalinated water will evaporate to the permeate (i.e., F) by the
time this temperature is reduced to 20 °C (assuming t = 0.3). This is typically
managed by reheating the cool brine reject and sending it back to the feed. In
DCMD, this recirculation is likewise done on the permeate side. Both pumps will now
be larger, by at least an order of magnitude, in order to achieve useful recoveries
exceeding 50%.

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• Electrical energy constraints: The thermodynamics of the simple MD setup in turn
constrains the electrical consumption. Each pump in Figure 6 will consume electrical
energy per unit water permeated, Eelec,std (kWh/m3
), according to:
Eelec,std = PF/ (η x F) x 1/3600 (10)
where PF is the MD module feed pressure (kPa), and η is pump efficiency. If we
assume PF = 20 kPa, and pump efficiency of 0.6, each pump consumes 0.12 kWh/m3
of electricity. Both pumps consume 0.23 kWh/m3
. Clearly achieving low pressure
drops along the module will have an impact on the electrical energy requirement of
MD systems. This minimum is related to the point above, where F equates to around
7.7 wt %;
• Thermal energy constraints: Water evaporation energy per unit mass, ΔHvap, is
2260 kJ/kg, or 628 kWh/m3
. This energy is in the form of thermal energy, which is
the standard thermal energy required to operate the MD system in Figure 6. This
value equates to a performance ratio (PR), or gain output ratio (GOR) of 1, being the
mass ratio of water produced to the amount of steam energy (i.e., latent heat) fed to
the process.
With state-of-the-art reverse osmosis requiring as little as 2 kWh/m3
of electric
energy and no thermal energy, we see that standard MD by thermodynamics uses an
order of magnitude less electricity, and nearly 300 fold the thermal energy to
desalinate the same amount of water. State-of-the-art MD systems feature
refinement of the system proposed in Figure 6, or its variants VMD, SGMD and
AGMD, primarily to reduce the thermal energy required, and more recently, the
electrical energy.
7.2. State of the Art MD Research and Systems
The principal research activities on MD can be divided broadly into two categories:
fouling/performance testing, and energy efficient process design. With
fouling/performance design, fundamental understandings of the diffusion
mechanisms coupled with heat and mass transfer has unlocked the critical science
needed to select optimal operating conditions, membrane materials and module
designs that ultimately give better flux performance for the same operational
conditions. Fouling of membranes has explored scaling issues for the classic
applications in brine concentration, and the more novel application in dairy
processing. While this research progresses to uncover further fundamental
improvements, the focus here is on the novel process configurations that address
the performance limitations defined in Section 2.1. The most notable organisations
specialising in MD modules or high efficiency systems are:

20
• Fraunhofer ISE (AGMD);
• Memstill and Aquastill (AGMD);
• Scarab (AGMD);
• Memsys (vacuum enhanced multi effect AGMD).
8. Mechanisms

21
8.1 Mass transfer
8.1.1 Direct Contact Membrane Distillation (DCMD)
Mass transfer in the DCMD process includes three steps: firstly the hot feed
vaporizes from the liquid/gas interface, secondly the vapour is driven by the vapour
pressure difference and crosses from the hot interface to the cold interface through
the pores, and thirdly the vapour condenses into the cold side stream. Therefore,
there are two major factors controlling the mass transfer: one is the vapour pressure
difference, and the other is the permeability of the membrane.
If the fluid dynamics conditions on both sides of the membrane could be considered
good, mass transfer through the membrane may be the limiting step for mass
transfer in MD. The influence of the physical properties on membrane permeability
includes:
(1) The effective area for mass transfer is less than the total membrane area because
the membrane is not 100% porous;
(2) For most practical membranes, the membrane pores do not go straight through
the membrane and the path for vapour transport is greater than the thickness of the
membrane; and
(3) The inside walls of the pores increase the resistance to diffusion by decreasing
the momentum of the vapour molecules.
The mass transport mechanism in the membrane pores is governed by three basic
mechanisms known as Knudsen-diffusion (K), Poiseuille-flow (P) and Molecular-
diffusion (M) or a combination between these known as the transition mechanism.
 Knudsen diffusion takes place when the pore size is too small, so the
collision between the molecules and the inside walls of the membrane
suitably expresses the mass transport and the collision between molecules
can be ignored.
 Molecular diffusion occurs when the molecules move corresponding to each
other under the influence of concentration gradients.
 In Poiseuille flow (viscous flow), the gas molecules act as a continuous fluid
driven by a pressure gradient.
The Knudsen number (Kn) is used to indicate the dominant mass transfer
mechanism in the pores:

22
The Knudsen number (Kn), defined as the ratio of the mean free path (λ) of
transported molecules to the membrane pore size, provides a guideline of which
mechanism is active inside the membrane pore. According to kinetic theory of gases,
the molecules are assumed to be hard spheres with diameter de and are involved in
binary collisions only. It is worth noting that the collision diameter for water vapour
and air are about 2.64×10−10
, and 3.66×10−10
, respectively. The average distance
travelled by molecules to make collisions (λ) is defined as.
λ = (kΒ x T) / ( x π x P x de
2
) (11)
kB, T and P are Boltzman constant, absolute temperature, and average pressure
within the membrane pores respectively. The mean free path value of water vapour
at 60 °C was estimated to be 0.11 μm.
 For Kn > 1 or dp < λ (Knudsen region), the mean free path of water vapour
molecules is large compared to the membrane pore size, which means the
molecule-pore wall collisions are dominant over molecule-molecule collision.
The mass transfer is:
CKn = 2π/3 x 1/ (R x T) x [(8 x R x T)/ (π x Mw)]1/2
x r3
/ (τ x δ) (12)
where ε, τ, r, δ and Mw are porosity, pore tortuosity, pore radius, membrane
thickness and molecular weight of water vapour, respectively.
 If the kn < 0.01 or dp > 100λ (continuum region), ordinary molecular diffusion
model represents the diffusion of the vapour flux through stationary air film
(the air which exist inside the membrane pores), ordinary molecular diffusion
is used to describe the mass transport
CD = π/ (R x T) x [(P x D)/ Pair] x r2
/ (τ x δ) (13)
where Pair is the air pressure within the membrane pore, D is diffusion coefficient,
and P is the total pressure inside the pore which is equal to the partial pressure of air
and water vapour.

23
In addition, the flux of water vapour molecules, which diffuse through the
membrane pores (stagnant air), is:
J = 1/ Pair x ε/ (τ x δ) x [(D x P x Mw)/ (R x T)] x ΔP (14)
where Pair and P are the average air pressure and average gas pressure within the
membrane respectively.
Removing the stagnant air existing inside the pores by degassing the feed and
permeate will reduce the molecular diffusion resistance, so the membrane
permeability will increase.
 However, If 0.01 < kn < 1 or λ < dp <100λ (transition region), the water vapour
molecules collide with each other, and also diffuse through the air film.
Consequently, the mass transfer takes place by both the Knudsen/ordinary
diffusion mechanism, where:
Cc = π/ (R x T) x 1/ (τ x δ) x [(2/3 x ((8 x R x T)/ (π x Mw)1/2
x r3
)-1
+ ((PD)/
Pα x r2
)-1
]-1
(15)
The diffusivity of water vapour through the stagnant air inside the pores is given by
PD = 1.895 x 10-5
x T2,072
(16)
In addition, the Fuller equation, which is a common equation to predict binary gas
diffusion, can be used
D = 10-7 x T1,75
x (1/ Mwα + 1/ Μwb)1/2
/ [Px ((Σvα)1/3
+ (Σvb)1/3
)2
] (17)
where Σv represents the diffusion volume, T is temperature in Kelvin and P is
pressure in atmospheres. The diffusion volume of air and water are 20.1 and 12.7
respectively.
It is stated that the molecule-pore wall collisions (Knudsen diffusion) and molecule-
molecule collisions (molecular diffusion) takes place simultaneously for pore size less
than 0.5 μm. Moreover, the flux can be expressed by molecular diffusion only for
large pores. Furthermore, the vapour flux across the membrane can be expressed by
Knudsen diffusion and Poiseuille (viscous) flow model for de-aerated DCMD. On the
other hand, the Poiseuille flow should be considered as one of the mechanisms of
mass transfer model in large pore size membrane.
It is helpful to analyze the transport in terms of resistances, to identify the
controlling role of each transport step, and as a result the flux permeate can be
improved. Table 4 shows DCMD membrane coefficients as reported by some
researchers.

24
It is believed that when the pore size is near the mean free path value (critical pore
size), the permeate flux under the Knudsen mechanism is higher than that obtained
from the combination of Knudsen and molecular diffusion mechanisms. Therefore,
choosing membranes that have small pore size may be better than membranes
having large pore size. It is worth mentioning that the effect of pore size distribution
can be neglected for large pore size.
8.1.2. Air Gap Membrane Distillation (AGMD)
The molecular diffusion theory is used to describe the transfer of vapour molecules
through the membrane and the air gap. A stagnant gas film (air) is assumed to lie
inside the membrane at the air gap side.
Kurokawa computed the flux by considering the diffusion in one direction through
both membrane and air gap, where the air gap is below 5 mm:
J = (P x Mw)/ (RT x P*) x (D/ (δ/ε3,6
+ 1) x ΔP (18)
where ΔP is the water vapour pressure difference between the feed on the
membrane and the condensation surface, and P* is the partial pressure of water.
Liu estimated the permeate flux for aqueous solution when the average operating
temperature, Ta, was between 30 °C and 80 °C, thus:
J = (Tf - Tp)/ (αΤα
-2,1
+ β) (19)
where α and β are parameters that can be determined experimentally.
It is worthwhile stating that the air gap is about 10 to 100 times the membrane
thickness, so the effect of air inside the membrane can be neglected.

25
Stefan diffusion was used to describe the diffusion through a stagnant gas film. It can
be represented mathematically as
Ν = - (c x D)/ (1 - y) x dy/ dz (20)
where D, y, c and z are diffusion coefficient, mole fraction of the vapour phase,
molar concentration and diffusion length, respectively.
The Stefan equation was solved by Kimura and Nakad
N = (c x D)/ z x ln((1-yf)/ ((1 - ym)) (21)
where ym and yf represent the mole fraction of vapour at the membrane and the
condensation film, respectively.
However, Jonson solved the same equation by neglecting the effect of temperature
and concentration polarization. They suggested that, the value of c D for water
vapour and air at around 40 °C to be calculated using this equation:
c x D = 6,3 x 10-5 x (22)
In addition, the molar concentration can be calculated from ideal gas law:
c = P/ (R x T) (23)
According to the standard condition, the diffusion coefficient can be corrected to the
desired temperature by:
D/ D0 = (T/ T0)3/2
(24)
Bouguecha used Stefan diffusion to express the vapour flux when it is governed by
diffusion through the membrane pores and by natural convection through the air
gap:
N = KT/ R x (Pf,m - Pfilm) (25)
where KT is the overall mass transfer coefficient. Stefan diffusion was also utilized to
evaluate the molar flux of seawater as:
N = DP/ (RT x l x Plm) x (P2- P4) (26)
where, P2, P4, Dw and Plm are the vapour pressures at Tf, m, the vapour pressures at
Tfilm, diffusion coefficient and log mean partial pressure respectively. The log mean
partial pressure difference at the air gap is defined as:
Plm = (P4 - P2)/ ln(P4/ P2) (27)

26
For a multi-component mixture, the Stefan-Maxwell equation was applied by Gostoli
and Sarti to express the ethanol and water vapour diffusion in stagnant gas (air). This
was given by:
dyi/ dz = / cDij x (yiNj - yjNi) (28)
The vapour composition at evaporation and condensation interface scan be
calculated by assuming liquid-vapour equilibrium, such that:
yi = (xi x ai x P0)/ P (29)
Vapour pressure P0 can be computed by the Antoine equation at the temperature of
interest. The activity coefficient ai can be calculated by the Van Laar equation at the
temperature and composition of interest. The condensate composition xi is
determined by the components flux
xi = Ni/ ΣN (30)
On the other hand, Banat and Simandl employed Stefan diffusion (Eq. (24)) to
represent the molar diffusion flux of an ethanol-water solution. The molar diffusion
flux of ethanol and water through stagnant gas (air) in terms of pressure is given by:
Ni = (ε x Di x P)/ (RT x l x Plm) x (Pi2-Pi4) (31)
For the non-equilibrium thermodynamics case, the ordinary diffusion, which is
related to the concentration gradient, and thermal diffusion which is related to the
temperature gradient were considered to calculate the total mass flux. A linear
relation between flux and vapour pressure can be assumed, and the thermal
diffusion can be neglected.
The Stefan–Maxwell model is reported to be more accurate than the molecular
diffusion model (Fick's law) for separation of azeotropic mixtures
8.1.3. Vacuum Membrane Distillation (VMD)
In order to remove air trapped in the membrane pores, the deaeration of the feed
solution or a continuous vacuum in the permeate side should be applied.
Consequently, the ordinary molecular diffusion resistance is neglected. The Knudsen
mechanism is used to express the mass transfer, Poisseille flow or both together.
 For example, when the ratio of the pore radius to the mean free path r/λ is
<0.05, the molecule-pore wall collisions control the gas transport mechanism
(Knudsen flow model) and the molar flow rate is:
Ni = 2π/3 x 1/RT x ((8RT/ (π x Mwi)1/2
x r3
/ (δ x τ) x ΔPi (32)

27
 If r is between 0.05λ and 50λ, both molecular-molecular and molecular-wall
collisions should be considered. The total mass transfer is described by the
Knudsen-viscous model and can be represented by the following equation:
Ni = π/ (RT x δ x τ) x (2/3 x (8RT/ (π x Mwi)1/2 x r3 + r4/ 8μi x Pavg) x ΔPi
(33)
where μi is the viscosity of species i, and Pavg is the average pressure in the pore.
 When r/λ is >50, molecular- molecular collision dominates and the mass
transfer can be expressed by Poisseuille flow (viscous), such that:
Ni = (π x r4
)/ 8μi x Pavg/ RT x 1/ (τ x δ) x ΔPi (34)
8.1.4. Sweeping Gas Membrane Distillation (SGMD)
The equations, which illustrate the mass transfer of DCMD can be used in SGMD.
Knudsen/molecular diffusion can be used to describe the mass transfer through the
membrane pores. Moreover, the circulation velocity and feed temperature are
significant parameters.
Sherwood correlation can be used to estimate the mass transfer coefficient, k, across
the boundary layers, then the concentration at the boundary layer can be evaluated.
The empirical form of the Sherwood correlation is
Sh = (k x d)/ D = (Constant) x Rea
x Scb
(35)
where Re, Sc, and D are Reynolds number, Schmidt number and diffusion coefficient
respectively (Table 5).
Schmidt numbers can be calculated by:

28
Sc= μ/ (ρ x d) (36)
where μ is the viscosity. For a non-circular channel, these correlations
can be utilized if the equivalent (hydraulic) diameter deq is employed.
deq = 4rH = 4S/ LP (37)
where rH, S and LP are the hydraulic radius, cross sectional area of the flow channel,
and length of wetted perimeter of the flow channel, respectively.
8.2. Heat transfer
In MD processes, heat and mass transfers are coupled together in the same direction
from the hot side to the cold side. Figure 4 illustrates these processes in DCMD,
which is typical for MD configurations. The feed temperature, Tf, drops across the
feed side boundary layer to T1 at the membrane surface. Some water evaporates and
is transported through the membrane. Simultaneously, heat is conducted through
the membrane to the cold (permeate) side. The cold flow temperature Tp increases
across the permeate boundary layer to T2 at the membrane surface on the cold side
as water vapour condenses into the fresh water stream and gains heat from the feed
side. The driving force is, therefore, the vapour pressure difference between T1 and
T2, which is less than the vapour pressure difference between Tf and Tp. This
phenomenon is called temperature polarization.

29
So two main heat transfer mechanisms occur in the MD system: latent heat and
conduction heat transfer and the heat transfer, which occurs in DCMD, can be
divided into three regions (Fig. 7)
Heat transfer by convection in the feed boundary layer:
Qf = hf x (Tf - Tf,m) (38)
Heat transfer through the membrane by conduction, and by movement of vapour
across the membrane (latent heat of vaporization). The influence of mass transfer on
the heat transfer can be ignored
Qm = Km/δ x (Tf,m - Tp,m) + J x ΔHv (39)
Qm = hm x (Tf,m - Tp,m) + J x ΔHv (43)
where hm represents the heat transfer coefficient of the membrane.
It is worth mentioning that hm can be rewritten for pure water or very diluted
solution, and where temperature difference across the membrane surfaces is less
than or equal to 10 °C by substituting Eq. J = Cm x dP/ dT x (Tf,m - Tp,m) into (39),
Qm = Km/δ x (Tf,m - Tp,m) + [Cm x dP/dT x (Tf,m - Tp,m)] x ΔHv (44)
Qm = [Km/δ + (Cm x dP/dT) x ΔHv] x (Tf,m - Tp,m)] (45)
Qm = hm x (Tf,m - Tp,m) (46)
For a non-linear temperature distribution assumption, Qm (for the x-dimension) is
also expressed as
Qm = -km x dT/dx + J x ΔHv (47)

30
For the permeate side, the convection heat transfer takes place in the permeate
boundary layer
Qp = hp x (Tp,m - Tp) (48)
At steady state, the overall heat transfer flux through the membrane is given by:
Q = Qf = Qm = Qp (49)
hf x (Tf - Tf,m) = km/δ x (Tf,m - Tp,m) + J x ΔHv = hp x (Tp,m - Tp) (50)
Q = U x (Tf - Tp) (51)
where U represents the overall heat transfer coefficient.
It is worth pointing out that the heat conduction can be neglected for non-sported
thin membrane and for high operating temperature as well. Moreover, the heat
transfer by convection is ignored in the MD process, except in AGMD.
The surface temperature of both sides of membrane cannot be measured
experimentally, or calculated directly. Therefore, a mathematical iterative model has
been designed to estimate these temperatures:
Tf,m = Tf - (J x ΔHv + km/δm x (Tf,m - Tp,m))/ hf (52)
Tp,m = Tp - (J x ΔHv + km/δm x (Tf,m - Tp,m))/ hp (53)
The value of Hv should be evaluated at average membrane temperature. However,
the model was evaluated at logarithmic average membrane temperature.
The surface membrane temperature in terms of temperature polarisation
coefficient, ψ, for pure water and very diluted solution,
Tf,m - Tp,m = 1/ (1 + H/ hf +H/ hp) x (Tf- Tp) = ψ x (Tf - Tp) (54)
where hm is equal to
hm = (Cm x dP/dT) x ΔΗv + km/δ (55)
The (Tf,m - Tp,m) is about 0.1 °C at low flux and does not exceed 0.5 °C at high flux.
Concerning the presence of free and force convection in laminar flow in DCMD, the
following equation to calculate the heat transfer coefficient is suggested:
Νu = 0,74 x Re0,2
x (Gr x Pr)0,1
x Pr0,2
(56)

31
For the AGMD configuration, the heat transfer through the AGMD could be
represented as in DCMD, except for the heat transfer across the air gap, which
occurs by conduction and vapour (mass transfer)
hf x (Tf - Tf,m) = J x ΔHv + km/δ x (Tf,m - Tp,m) = J x ΔHv + kg/l x (p,m - Tfilm) = hd
x (Tfilm - T5)(57)
In addition, we can use the following equation to calculate the heat transfer
coefficient for the condensate film (pure vapour) on a vertical wall:
hd = 2/3 x x [(kfilm
3
x ρ2
x g x ΔHv)/ (μ x L x (Tfilm - T5)]1/4
(58)
The sensible heat for the MD system can be neglected, because it has a very small
magnitude compared to the heat of vaporization
Q = J x ΔHv (59)
Free convection heat transfer between two vertical plates is also used to describe
the heat transfer phenomenon in the air gap region, when the air gap distance is
over 5 mm
Nu = c x (Pr x Gr)n
x (l/L)1/9
(60)
where
105
< Gr < 107
, c = 0,07 and n = 1/3
104
< Gr < 105
, c = 0,2 and n = 1/4
For VMD configuration, heat transfer by convection in the feed boundary layer can
be expressed as:
Qf = hf x (Tf - Tf,m) (61)
However, the heat transfer by conduction through the membrane is ignored, so the
heat transfer across the VMD can be written as:
hf x (Tf - Tf,m) = J x ΔHv (62)
For SGMD, the heat transfer equations, which describe the DCMD can be used.
The heat transfer coefficients of the boundary layers can be estimated by the Nusselt
correlation (see Table 6). Its empirical form is:
Nu = Constant x Rea
x Prb
(63)

32
Consequently, the heat transfer coefficient h can be calculated using Reynolds and
Prandtl numbers(Re and Pr), i.e.
Reynolds number = Re = (ν x d x ρ)/ μ (64)
Prandtl number = Pr = (cp x μ)/ k (65)
Grashoff number = Gr = (g x β x ΔΤ x L3
x ρ2
)/ μ3
(66)
where v, ρ, μ, cp, g, β, L and k are fluid velocity, density, viscosity, heat capacity,
gravity acceleration, thermal expansion coefficient, height and thermal conductivity.
The mass transfer and the heat transfer can be related, by:
Sh x Sc-1/3
= Nu x Pr-1/3
(67)

33
9. Thermal efficiency and energy consumption
The thermal efficiency Π in MD can be specified as the ratio of latent heat of
vaporization to the total (latent and conduction) heat. The thermal efficiency can be
improved by increasing the feed temperature, feed flow rate and membrane
thickness. In contrast, it decreases when the concentration for salt solution
increases.
For DCMD, the thermal efficiency Π can be expressed as:
Π = (J x ΔΗv)/ (J x ΔHv + km/δ x (Tf,m - Tp,m)) (68)
For pure water, the characteristics of the membrane, such as porosity and tortuosity,
determine the thermal efficiency, with no dependence on membrane thickness.
Around 50–80% of the total heat flux across the membrane is considered to be
latent heat; whereas 20–40% of heat is lost by conduction through the membrane.
The heat lost by mass flux can be estimated by:
Qlost/J = km/Cm x (Tf,m - Tp,m)/ (P2-P3) (69)
For a very dilute solution, and low membrane temperature we can use the following:
Qlost/J = km/Cm x 1/ (dP/dT)Tm (70)
Working at a high temperature and flow rate reduces the heat loss.
There are three forms for heat transfer to be lost in the DCMD system. The first form
is due to the presence of air within the membrane. Secondly, heat loss through the
membrane by conduction, and finally by temperature polarization. Solutions to
minimize heat loss in the DCMD, are: de-aeration of the feed solution, increasing the
membrane thickness, creating an air gap between the membrane and the
condensation surface, and operating within a turbulent flow regime.
In terms of AGMD thermal efficiency, suggested that the thermal efficiency is
proportional to the membrane distillation temperature difference. They introduced
two parameters α and β, which can be determined experimentally for an air gap less
than 5 mm, and average membrane distillation temperature, Ta, between 30 °C and
80 °C by:
η = 1 - (α x Tα
-2.1
)/λ x ((Tf-Tp) x Cp)/( α x Tα
-2.1
+ β) + kα/l) (71)
cp and ka are specific heat and air gap thermal conductivity.

34
It is observed that by increasing the feed temperature from 40 °C to 80 °C, the
thermal efficiency increased by 12%, whereas the salt concentration has a marginal
effect on the thermal efficiency.
With regard to energy consumption, if we use a simple energy balance to compute
the energy consumption of hot and cold streams for DCMD and VMD using different
flow configurations,
Q = m x cp x ΔΤ (72)
We can found that the cross-flow configuration is the best, in terms of high flux and
energy consumption. Moreover, hybrid RO/MD becomes the best choice when an
external energy source is available. In addition, heat transfer to the cooling side by
heat conduction, and by heat of condensation can be used (recovered) to preheat
the feed solution, which minimizes the heat requirement and improves the
operation cost. The percentage of heat recovery depends on the heat exchanger
area. It is pointed out that the MD performance rises by 8% when heat recovery is
used. The heat exchanger capacity should be optimized with membrane area, in
order to get high production flux for a solar powered membrane distillation system.
From the economic point of view, the capital cost is very sensitive to heat recovery,
because the heat exchanger is the most expensive item in a solar-powered MD plant.
We must optimized the solar collector area, membrane area and heat recovery to
achieve low capital cost and high flux.
10. Temperature polarization and concentration polarization
Since the vaporization phenomenon occurs at the membrane hot surface and
condensation at the other side of membrane, thermal boundary layers are
established on both sides of the membrane. The temperature difference between
the liquid-vapour interface and the bulk is called temperature polarization, ψ , which
is defined as:
ψ = (Tf,m - Tp,m)/ (Tf - Tp) (73)
Lawson represented ψ with slight difference for VMD as:
ψ = (Tf - Tm,f)/ (Tf - Tp) (74)
The effect of heat transfer boundary layer to total heat transfer resistance of the
system is measured by temperature polarization.
When the thermal boundary layer resistances are reduced, the temperature
difference between the liquid-vapour interface and the bulk temperature becomes
close to each other and, consequently, ψ approaches 1, which means a typical

35
system. On the other side, zero ψ means a high degree of concentration polarization
is taking place, and the system is controlled by large boundary layer resistance.
Usually, the value of ψ lies between 0.4–0.7 for DCMD. It is pointed out that
temperature polarization becomes important at high concentration, high
temperature and low feed velocity.
Concentration polarization, Φ is defined as the increase of solute concentration on
the membrane surface (cm) to the bulk solute concentration(cf):
Φ = cm /cf (75)
In order to estimate the concentration of the solute (mole fraction) on the
membrane surface, the following relation is suggested:
cm = cf x exp(j/(ρ x K)) (76)
where ρ is the liquid density and K is mass transfer coefficient.
Concerning the influence of high concentration on mass transfer coefficient and
distilled flux, the viscosity, density of the feed, solute diffusion coefficient, and the
convective heat transfer coefficient are directly related to the concentration and
temperature. The concentration polarization and fouling must be considered in
modelling, and the permeate flux cannot be predicted by Knudsen, molecular and
Poiseuille flow, because the properties of the boundary layer at the membrane
surface vary from the bulk solution.
11. Fouling
Membrane fouling is a major obstacle in the application of membrane technologies,
as it causes flux to decline. The foulant, e.g., bio-film, precipitations of organic and
inorganic matter, can reduce the permeability of a membrane by clogging the
membrane surface and/or pores. Although membrane distillation is more resistant
to fouling than conventional thermal processes, dosing of anti-scalants can be used
to control scaling. Lower feed temperatures can substantially reduce the influence of
fouling in DCMD.
Since the hydrophobic MD membrane is the barrier between the feed and permeate,
membrane wetting will reduce the rejection of the non-volatiles. Membrane wetting
can occur under the following conditions:
• The hydraulic pressure applied on the surface of the membrane is greater than the
LEP;
• The foulant depositing on the membrane surface can effectively reduce the
hydrophobicity of the membrane, which was generally found in a long-term

36
operation or in treating high-concentration feeds such as for brine crystallisation;
and
• In the presence of high organic content or surfactant in the feed, which can lower
the surface tension of feed solution and/or reduce the hydrophobicity of the
membrane via adsorption and lead to membrane wetting.
The fouling problem is significantly lower than that encountered in conventional
pressure-driven membrane separation. It is pointed out that membrane fouling by
inorganic salt depends on the membrane properties, module geometry, feed
solution characteristic and operating conditions. There are several types of fouling,
which may block the membrane pores. Biological fouling is growth on the surface of
the membrane (by bacteria), and scaling (for the high concentration solution), which
will create an additional layer on the membrane surface, composed of the particles
present in the liquid.
Fouling and scaling lead to blocking the membrane pores, which reduces the
effective membrane, and therefore the permeate flux obviously decreases. These
may also cause a pressure drop, and higher temperature polarization effect. The
deposits formed on the membrane surface leads to the adjacent pores being filled
with feed solution (partial membrane wetting). Moreover, additional thermal
resistance will be created by the fouling layer, which is deposited on the membrane
surface. As a result, the overall heat transfer coefficient is changed. For DCMD at
steady state:
hfx(Tf - Tf,fouling) = kfouling/δfoulingx(Tf,fouling - Tf,m) = km/δx(Tf,m - Tp,m) + JxΔHv =
hpx(Tp,m - Tp) (77)
where kfouling, δfouling and Tf, fouling are the fouling layer thermal conductivity, thickness,
and fouling layer temperature, respectively.
Concerning the effect of high concentration of NaCl and Na2SO4 on the permeate
flux, the flux gradually decreases during the MD process, until the feed
concentration reaches the super saturation point, and then the flux decrease sharply
to zero. Afterwards, the membrane was completely covered by crystal deposits.
When the membrane surface concentration reaches saturation, the properties of the
boundary layer will differ from the bulk solution properties. Currently, pre-treatment
and membrane cleaning are the main techniques to control fouling. Pre-treatment
process increases the product flux by 25%, which means that the pre-treatment
process is important, in order to enhance the permeate flux. Fouling intensity can be
limited by operating at low temperature (feed temperature), and increasing the feed
flow rate.

37
Notes: θf is the angle
between spacer fibres in
the flow direction; lm is the
distance between parallel
spacer fibres; hsp is the
height of the spacer and df
is the diameter of a single
spacer fibre.
12. Operating parameters
In this section, the influence of feed temperature, concentration and air gap will be
reviewed and major findings will be cited and discussed.
12.1 Parameters to Reducing Temperature Polarization
To maximise flux, it is necessary to increase the vapor pressure difference across the
membrane or to reduce temperature polarization. Therefore, it is necessary to
improve the convective heat transfer coefficient for the purpose of producing more
flux. The convective heat transfer coefficient can be expressed as:
αf = - λf/ (Tf - T1) x (dT/dy)boundary (78)
where λf is thermal conductivity of the feed, and (dT/dy)boundary is the temperature
gradient in the thermal boundary layer of the feed. From Equation (10b), it can be
seen that the convective heat transfer coefficient can be improved effectively by
reducing the thickness of the thermal boundary layer. As the thickness of the
thermal boundary layer can be reduced by enhancing the stream turbulence,
increasing flow rate can effectively improve the flux. However, the hydrodynamic
pressure has a square relationship to the flow rate, and the increased pressure will
diminish the effect of increasing turbulence if the membrane is compressible.
The presence of turbulence promoters, e.g., net-like spacers or zigzag spacers shown
schematically in Figure 5 can effectively reduce the thickness of the thermal
boundary layer and improve αf. It is also important that high heat transfer rates are
achieved with a low pressure drop in the channels where the feed solution and
cooling liquid are flowing.

38
From reported data, it is found that the temperature polarization coefficient of
spacer-filled channels falls in the range of 0,9–0,97, in comparison with a
temperature polarization coefficient 0,57–0,76 for flowing channels without spacers.
It is also noticed that the influence of turbulence on flux becomes less at higher
turbulence levels. Therefore, it is necessary to control turbulence within an adequate
range to reduce the energy cost associated with pumping.
12.2 Feed temperature
As MD is driven by vapor pressures which vary exponentially with the stream
temperature, the flux is affected greatly by the feed temperature. Furthermore,
since the heat loss through thermal conduction is linear to the temperature
difference across the membrane as according to Equation (3), the proportion of
energy used for evaporation will increase as the feed temperature increases.
However, an increase of temperature polarization due to the high flux and greater
heat and mass transfer was also observed with rising temperature, but this can be
reduced by using turbulence promotors such as spacers.
As can be seen in Table 7, the feed temperature has a strong influence on the
distilled flux. According to the Antoine equation, the vapour pressure increases
exponentially with temperature. Therefore, the operating temperature has an
exponential effect on the permeate flux.
At constant temperature difference between the hot and the cold fluid, the
permeate flux increases when the temperature of the hot fluid rises, which means
the permeate flux is more dependent of the hot fluid temperature. It is pointed out
that increasing the temperature gradient between the membrane surfaces will affect
the diffusion coefficient positively, which leads to increased vapour flux. Similarly, it
is believed that there is a direct relation between diffusivity and temperature, so
that working at high temperature will increase the mass transfer coefficient across
the membrane. Moreover, temperature polarization decreases with increasing feed
temperature. In terms of coolant temperature, a noticeable change takes place in

39
the permeate flux when the cold side temperature decreases. In addition, more than
double permeate flux can be achieved compared to a solution, at the same
temperature difference. However it was found that the effect of the cold side
temperature on the permeate flux is neglected at fixed hot side temperature,
because of low variation of vapour pressure at low temperatures.
12.3 The concentration and solution feature
There is a significant fall in the flux product when feed concentration increases due
to decreasing vapour pressure and increasing temperature polarization and the
reduction in product flux is linear with time. Furthermore, there is a reduction in the
permeate when the acid concentration increase. About 12% reduction in permeate
flux happened when the feed (NaCl) increased from 0 to 2 Molar concentration. This
decrease in the permeate flux amount is due to the reduction in the water vapour
pressure. Lawson and Lloyd studied the reasons for decreasing product flux when
the concentration of NaCl increases.
They found three reasons for this reduction;
1) water activity, which is a function of temperature, decreases when the
concentration increases
2) the mass transfer coefficient of the boundary layer at the feed side decreases due
to increased influence of concentration polarization, and
3) the heat transfer coefficient decreases as well at the boundary layer, because of
the reduction in the surface membrane temperature.
Therefore, the vapour pressure of the feed declines, which leads to reduced
performance of MD. The viscosity is an important factor in flux reduction. The heat
transfer coefficient decreases due to the reduced Reynolds number. The effect of
thermal conductivity and heat capacity on the flux reduction is negligible.
Furthermore, the impact of density on the flux production is important for salt
solutions. There is variation in the permeate flux with time (see Table 8), and that it
is difficult to calculate the permeate flux using existing models.

40
12.4 Recirculation rate
Table 9 summarizes the effect of recirculation rate. Working at a high recirculation
rate minimizes the boundary layer resistance and maximizes the heat transfer
coefficient.
As a result, higher flux can be achieved. It is indicated that the increasing volumetric
flow rate will enhance the permeate flux. The fluid velocity rises when the
volumetric flow rates increases, so that the convective heat transfer coefficient
develops and the thermal boundary layer thickness decreases. As a result, the
temperature polarization effect reduces. Moreover, there is a significant change in
temperature polarization, when the rate of recirculation changes. This is because the
recirculation rate enhances the heat transfer, which leads to rise in the product flux
and temperature polarization. The effect of the cold side flow rate on the permeate
flux has yet to be decided.
12.5 The air gap
It is suggested that the flux declines linearly with 1/l . Reducing the air gap width will
increase the temperature gradient within the gap, which leads to increased
permeate flux. Table 10 summarizes the air gap effect on the permeate flux.

41
12.6 Membrane type
The membrane permeation flux is proportional to the porosity, and inversely
proportional to the membrane thickness and tortuosity. For a larger pore size
membrane, higher permeate flux is obtained. In addition, higher flux is achieved
using a membrane without support, compared to the same membrane pore size
with support. For a more efficient MD process, low thermal conductivity material
should be used (unsupported membrane).
13. Future Developments and Conclusions
MD has been explored since the early 1960s, but only in the last decade has the
interest grown substantially such that commercial systems are readily available,
backed with pilot trial experience. Various MD providers offer solutions that are
primarily focused on minimizing thermal energy demand, but there are also
possibilities to reduce electrical energy demand.
MD has been used mainly trailed for removing salt for sea water and brackish water.
It has also proven to be a suitable technology for removal of other contaminants,
such as heavy metals, radionuclides, and organics from brackish, produced, industrial
and other impaired water. While it is capable of treating many kinds of water, its
ability to compete with established technologies such as RO, ED, MED and MSF is
currently limited due to its high energy use. Consequently it is likely to find
application where current established technologies are unable to operate or in
applications that substantially favour its use. For instance, the treatment of brine
streams that reverse osmosis finds difficult to treat may be a possible application,
and integration of MD with RO to treat RO brine may be a suitable application where
brine disposal is problematic. Treatment of CSG water brine is one such potential
application, where reduction of brine pond areas has substantial capital cost
benefits.
Similarly, application in industries that have significant low grade waste heat
sources, such as power stations and chemical plants, would also seem to be strong
candidates for application of MD. The high quality of MD permeate compared to RO
permeate may also provide advantages in these applications, particularly if purified
water is required as boiler feed.
Finding suitable applications for MD currently seems to be the major impediment to
its wider commercial use. The theory of its operation is well known, and models are
available to allow design and scale up of MD systems using local heat sources. The
ability to design MD processes using site specific heat flows is critical for its
application, as it is dependent upon waste heat sources to achieve economic
advantages, and the quality and available heat flows from such heat sources will vary

42
from site to site. Efficient designs will be required to take this variability in available
heat in to account.
Low fluxes and wetting have also been limitations for MD implementation. Having
highly permeable membranes and suitable modules with improved hydrodynamics
will allow increased permeate flux and overall performance of the MD process.
Membrane hydrophobicity and pore geometry are critical parameters in reducing
MD membrane wetting, and surface coatings are enabling reduced wetting to be
achieved. For example, oleophobic coatings can reduce wetting and fouling from oily
feeds. Membrane hydrophobicity also determines the largest possible membrane
pore size for scale up as do process parameters such as feed water temperature,
operating pressure, flow rate, and liquid composition.
A large variety of materials has been tested and investigated as MD membranes. It
appears that although morphological features are critical to achieve high flux,
improving the membrane performance is a complex issue involving a number of
parameters. The variety of the MD configurations, membrane morphologies, module
shape and size, as well as the testing conditions cause the large scatter.
As defined by theory, pore size, porosity and thickness of the active layer matter,
and the characteristics of the support are critical to achieving high flux. Controlling
the thermal transfers across the different strata of the membrane is also critical and
more efforts should be focused on improving the interfaces between the active layer
and the supporting layer in order to reduce temperature polarization effects.
Thermal conductivity measurements are often difficult to perform due to the
difficulty in controlling the interfacial contact, but this should be an area of focus for
researchers in order to better understand their structures. Although a few studies
did investigate the long term performance of their membranes, it would also be
interesting to investigate the long term flux and rejection stability of these novel
membranes as very few groups investigated the impact of contaminants, such as
chlorine, or chemical and thermal degradation on the process. MD induces strong
temperature gradients across the membranes, and thermal degradation could occur
over time depending on the composition of the feed. In addition, the compressibility
of the membrane when stressed in the module under the pressure difference will
likely affect flux and energy efficiency, particularly for DCMD.
To date commercial large pore size PTFE flat sheet membranes still show higher
permeance than laboratory fabricated membranes when tested under similar
conditions. A number of routes are open for researchers to improve the
performance of membranes for MD. These routes include fabricating smaller pore
size, but thinner active membrane layers with more hydrophilic materials. The
smaller pore size will then lead to a larger LEP reducing the risk of liquid water
penetration into pores while hydrophilic surfaces may reduce fouling. Research

43
could also, on the other hand, be driven towards the processing of larger pore size
hydrophobic membranes to achieve higher water vapor permeability in order to
become more competitive with commercially available structures. Tuning the
surface energy of the membrane is also critical, and novel approaches combining
hydrophilic and hydrophobic materials have shown highly promising results. Other
routes include the control of the support morphology by introducing large macro
cavities to maximize the liquid water or vapor transport and reducing possible heat
and concentration polarization effects.
Ceramic membranes are possible candidates in place of polymeric membranes in MD
applications due to higher thermal resistance, mechanical strength, chemical
stability and oxidant tolerance. Additional research is required to find optimal
chemical modification candidates as well as optimal procedures to change the
hydrophilic inorganic membranes to hydrophobic membranes without compromising
the performance and permeate flux of the MD process. Nanoparticles are important
emerging candidates to be used in the manufacturing of membranes for MD. They
allow for control of membrane wetting and fouling. Graphene and carbon nanotubes
are the most promising candidates due to their physico-chemical properties, which
help engineering of desired structures and selectivity of the membrane separation
process. Electro-spun webs, which are manufactured as affinity membranes for the
study and growth of biological cells, may open opportunities for research in the area
of membranes for MD.
MD appears to be poised for commercial implementation, and identification of
opportunities that maximise the advantages of MD over competing technologies is
emerging. In developing these opportunities, the energy consumption and desalted
unit cost will decrease; therefore, competitive values with those of other
desalination processes can be reached.

44
Membrane Distillation (MD)
Reverse osmosis (RO)
Ultrafiltration (UF)
Direct Contact Membrane Distillation (DCMD)
Air Gap Membrane Distillation (AGMD)
Sweeping Gas Membrane Distillation (SGMD)
Thermostatic sweeping gas membrane distillation (TSGMD)
Vacuum Membrane Distillation (VMD)
Polytetrafluoroethylene (PTFE)
Polypropylene (PP)
Polyvinylidene fluoride (PVDF)
Liquid entry pressure (LEP)
Scanning Electron Microscopy (SEM)
Atomic Force Microscopy (AFM)
Direct Contact Membrane Distillation (DCMD)

An Introduction to Membrane Distillation

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