Reverse Osmosis and Membrane Distillation

Membrane distillation and treatment of
Reverse Osmosis reject
MSc Christos Charisiadis

Contents
1. Introduction to Membrane Distillation 1
2. Process fundamentals 2
2.1 Configurations of MD Modules 5
2.2 Membranes for Membrane Distillation Applications 6
2.3 Membrane Materials 7
2.4 Characteristics of MD Membrane 9
2.4.1 Contact angle 11
2.4.2 Liquid entry pressure and wetting phenomena 12
3. Transport mechanisms and polarization phenomena 12
3.1 Theory of heat transfer 12
3.2 Theory of mass transfer 14
4. Process parameters 17
4.1 Parameters to Reducing Temperature Polarization 19
5. Long-term performance; Membrane Fouling and Wetting 20
6. Engineering aspects; MD applications 22
7. Advances on MD Processes and Modules for Water Purification 24
7.1 MD Stand-Alone Systems 24
7.2 State of the Art MD Research and Systems 26
7.3 Hybrid MD Systems 26
7.3.1 MD Integration with RO or NF 26
7.3.2 MD Integration with FO 27
8. Brackish water 28
8.1 RO/ED/EDR Concentrate 29
8.2 Concentrate Management Cost for MD 30
9. Investigation of high recovery of concentrated RO brine using MD 32

9.1 Brine chemical analysis 33
9.2 Vacuum enhanced direct contact MD 34
9.2.1 Water flux and recovery 34
9.2.2 VEDCMD membrane cleaning 37
9.2.3 VEDCMD with scale inhibitor 38
9.2.4 VEDCMD water recovery 39
9.3 Comparing VEDCMD and FO for brine treatment 39
9.4 Conclusions 40
9.5 Recommendations 41
10. Membrane distillation as a means for reverse osmosis concentrate volume
minimization 42
10.1 Comparison of emerging technologies for concentrate treatment 43
11. PRO concentrate treatment with DCMD 47
12. RO concentrate treatment with VMD 54
12.1 Experimental 55
12.2 Results and discussion 56
12.3 VMD performance with concentrated synthetic brines 60
12.4 Observation and study of scaling 66
12.5 Study of scaling for SW300 solution, the highest feed concentration solution
68
12.6 Membrane distillation of actual RO brines 69
12.7 Conclusion 70
13. Integration of accelerated precipitation softening with MD for PRO concentrate
70
13.1. PRO concentrate and reagents 71
13.2 APS-DCMD process 72
13.3 Membrane module 72

13.4 Determination of the optimum softening conditions 73
13.4.1 Initial pH 73
13.4.2 Seed selection and dosage 74
13.5 Performance of integrated APS-DCMD process 75
13.6 Conclusions 77
14. Sustainable operation of MD for mineral recovery from hypersaline solutions
77
14.1 DCMD for concentration of super saturated solutions in mineral production
78
14.2 Materials and methods 78
14.2.1 Membranes 78
14.2.2 Bench-scale system 78
14.3 Results and discussion 80
14.3.1 Pure water permeability experiments 80
14.4 Direct contact membrane distillation batch experiments 80
14.4.1 Successive batch experiments: water flux and salt rejection 80
14.4.2 Membrane scaling investigation 83
14.4.3 Extended scaling experiments 85
14.5 Scaling mitigation techniques 86
14.5.1 Mitigating rapid flux decline 86
14.5.2 Flow reversal 87
14.5.3 Temperature reversal 88
14.6 Efficiency of MD over natural evaporation 89
14.7 Conclusions 90
15. References 91

1
MD and treatment of RO reject-Christos Charisiadis
1. Introduction to Membrane Distillation [1]
Membrane Distillation (MD) is one of the emerging non-isothermal membrane
separation processes known for about 47 years and still needs to be developed for
its adequate industrial implementation. It refers to a thermally driven transport of
vapour through non-wetted porous hydrophobic membranes, the driving force being
the vapour pressure difference between the two sides of the membrane pores.
As in other membrane separation processes, the driving force is the chemical
potential difference through the membrane thickness. Simultaneous heat and mass
transfer occur in this process and, as will be explained later, different MD
configurations such as
(i) direct contact membrane distillation,
(ii) sweeping gas membrane distillation,
(iii) vacuum membrane distillation and
(iv) air gap membrane distillation, can be used for various applications (desalination,
environmental/waste cleanup, water-reuse, food, medical, etc.)
The involved simultaneous heat and mass transfer phenomena through the
membrane, the different MD configurations and the various MD applications make
MD attractive within the academic community as a kind of didactic application.
Additionally, the possibility of using waste heat and/or alternative energy sources,
such as solar and geothermal energy, enables MD to be combined with other
processes in integrated systems, making it a more promising separation technique
for an industrial scale. Furthermore, the lower temperatures than in the
conventional distillation, the lower operating hydrostatic pressures than in the
pressure-driven processes (i.e., reverse osmosis (RO), nanofiltration (NF),
ultrafiltration (UF) and microfiltration (MF)), the less demanding membrane
mechanical properties and the high rejection factors achievable especially during
water treatment containing non-volatile solutes make MD more attractive than any
other popular separation processes.
Advantages of membrane distillation over reverse osmosis or other thermal methods
of desalination include [3]:
• It produces very high-quality distillate. In most circumstances salt rejections of 99-
100% are achievable.
• Water can be distilled at relatively low temperatures (i.e. 5 to 80°C). As the driving
force for MD is temperature difference, very low feed temperatures can produce

2
reasonably high rates of product water and may be more practical considering the
nature of some water impurities (e.g. scaling issues at high temperature).
• Low-grade heat such as industrial waste heat, solar or desalination waste heat may
be used.
• The feed water does not require extensive pre-treatment that is typically required
for pressure-based membrane processes.
Unfortunately, from the commercial stand point, MD has gained only little
acceptance and is yet to be implemented in the industry. The major barriers include
MD membrane and module design, membrane pore wetting, low permeate flow rate
and flux decay as well as uncertain energetic and economic costs.
2. Process fundamentals
MD is a thermally driven process, in which water vapour transport occurs through a
non-wetted porous hydrophobic membrane. The term MD comes from the similarity
between conventional distillation process and its membrane variant as both
technologies are based on the vapour-liquid equilibrium for separation and both of
them require the latent heat of evaporation for the phase change from liquid to
vapour which is achieved by heating the feed solution. The driving force for MD
process is given by the vapour pressure gradient which is generated by a
temperature difference across the membrane. As the driving force is not a pure
thermal driving force, membrane distillation can be held at a much lower
temperature than conventional thermal distillation. The hydrophobic nature of the
membrane prevents penetration of the pores by aqueous solutions due to surface
tensions, unless a transmembrane pressure higher than the membrane liquid entry
pressure (LEP) is applied. Therefore, liquid/vapour interfaces are formed at the
entrances of each pore. The water transport through the membrane can be
summarized in three steps:
(1) formation of a vapour gap at the hot feed solution–membrane interface;
(2) transport of the vapour phase through the microporous system;
(3) condensation of the vapour at the cold side membrane–permeate solution
interface.
Various MD configurations can be used to drive flux. The difference among these
configurations is the way in which the vapour is condensed in the permeate side.
Figure 1 illustrates the four commonly used configurations of MD described as
follows:

3
Fig.1, Schematic representation of MD configurations [2]
1. In direct contact membrane distillation (DCMD), water having lower temperature
than liquid in feed side is used as condensing fluid in permeate side. In this
configuration, the liquid in both sides of the membrane is in direct contact with the
hydrophobic microporous membrane. DCMD is the most commonly used
configuration due to its convenience to set up in laboratory. However, direct contact
of the membrane with the cooling side and poor conductivity of the polymeric
material results heat losses throughout the membrane. Therefore, in DCMD the
thermal efficiency which is defined as the fraction of heat energy used only for
evaporation, is relatively smaller than the other three configurations.
2. In air gap membrane distillation (AGMD), water vapour is condensed on a cold
surface that has been separated from the membrane via an air gap. The heat losses
are reduced in this configuration by addition of a stagnant air gap between
membrane and condensation surface.

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3. In sweeping gas membrane distillation (SGMD), a cold inert gas is used in
permeate side for sweeping and carrying the vapour molecules to outside the
membrane module where the condensation takes place. Despite the advantages of a
relatively low conductive heat loss with a reduced mass transfer resistance, due to
the operational costs of the external condensation system, SGMD is the least used
configuration.
4. In vacuum membrane distillation (VMD), the driving force is maintained by
applying vacuum at the permeate side. The applied vacuum pressure is lower than
the equilibrium vapour pressure. Therefore, condensation takes place outside of the
membrane module.
Each of the MD configurations has its own advantages and disadvantages for a given
application (Table 1),
Table 1, Advantages, disadvantages and application areas for MD configurations [5]
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 [4].

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Mainly, MD should be applied for non isothermal membrane operations in which the
driving force is the partial pressure gradient across the membrane that complies
with the following characteristics [1]:
i) Porous.
ii) Not wetted by the process liquids.
iii) Does not alter the vapour/liquid equilibrium (VLE) of the involved species.
iv) Does not permit condensation to occur inside its pores.
v) Is maintained in direct contact at least with the hot feed liquid solution to be
treated.
2.1 Configurations of MD Modules [4]
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.
In plate and frame modules, the membranes which are usually prepared as discs or
flat sheets are placed between two plates. The feed solution flows through flat,
rectangular channels. Polymeric flat sheet membranes are easy to prepare, handle,
and mount. The same module can be used to test many different types of MD
membranes. The membrane can be supported to enhance mechanical strength [2].
Tubular, capillary or hollow fiber membrane modules are shell and tube type
modules housing pressure-tight tubes. The support is not needed in this type of
modules. The membranes are usually a permanent integral part of the module and
are not easily replaced. Tubular membrane modules provide much higher membrane
surface area to module volume ratio than plate and frame modules. These modules
offer higher cross-flow velocities and large pressure drop and generally used for MD
of high viscous liquids. The production costs are very low and membrane fouling can
effectively be controlled by the proper feed flow and back-flushing of permeate in
certain time intervals. The main disadvantage of the capillary membrane module is
the requirement of low operating pressure (up to 4 bars) [2].
Fig.2, MD modules (a) Tubular module for hollow fiber, (b) Plate and frame module for flat sheet membrane

6
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
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.
2.2 Membranes for Membrane Distillation Applications [4]
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.
Hollow fiber module has the highest packing density of all module types. Its
production is very cost effective and hollow fiber membrane modules can be
operated at pressures in excess of 100 bars. The main disadvantage of the hollow
fiber membrane module is the difficult control of membrane fouling. Therefore, a
proper pretreatment should be applied for separation of macromolecules [2].

7
Fig.3, Schematics of (a) hollow fiber and (b) flate sheet membranes, [4]
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.
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.
As the flux from membrane distillation is related to the membrane length, it is more
appropriate to compare membrane performance with the mass transfer coefficient
rather than the flux. However, it is difficult to calculate the mass transfer coefficients
from published works, because typically there is insufficient provision of data.
Therefore, the flux provided here is only used as an approximate indication of
performance.
2.3 Membrane Materials [2]
The selection of the membrane is the most crucial factor in MD separation
performance. As stated earlier, the membrane used for MD process must be
hydrophobic and porous. There are various types of membranes meeting these
expectations; however the efficiency of a given MD application depends largely on
additional factors such as resistance to mass transfer, thermal stability, thermal
conductivity, wetting phenomena and module characterization [2].

8
A large variety of membranes including both polymeric and inorganic membranes of
hydrophobic nature can be used in MD process; however polymeric membranes
have attracted much more attention due to their possibility to modulate the intrinsic
properties. Polytetrafluoroethylene (PTFE), polypropylene (PP) and
polyvinylidenefluoride (PVDF) are the most commonly used polymeric membranes
due to their low surface tension values (Table 2) [2].
Table 2, Critical surface tension values of some polymers, [4]
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 3
[4].
Table 3, Reported surface energy and thermal conductivity of most popular materials used in MD [4]
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 (however this polymer easily dissolves at room temperature in
a variety of solvents including dimethylformamide (DMF) and triethylphosphate
(TEP) [2]). 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 [4].

9
2.4 Characteristics of MD Membrane [2]
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. Thus, the
properties of membranes suitable for membrane distillation should include [4]:
• 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;
• 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.
• 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;
• Low thermal conductivity. High thermal conductivities increases sensible heat
transfer and reduce vapor flux due to reduced interface temperature difference; and
• 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.
There are some additional criteria that should be taken into consideration for
selection of the appropriate membrane for a given MD application such as pore size,
tortuosity, porosity, membrane thickness and thermal conductivity. The relationship
between the transmembrane flux and the different membrane characteristic related
parameters is given by
N x α x (<rα
> x ε)/ (τ x δ) (1)
where Ν is the molar flux, <rα
> is the mean pore size of the membrane pores where
α equals 1 for Knudsen diffusion and equals 2 for viscous flux, ε is the membrane
porosity, τ is the membrane tortousity and δ is the membrane thickness.

10
Membrane pore size: Membranes with pore sizes ranging from 10 nm to l μm can be
used in MD. The permeate flux increases with the increase in pore size as
determined by Knudsen model. However, in order to avoid wettability, small pore
size should be chosen. Thus, an optimum value for pore size has to be determined
for each MD application depending on the type of the feed solution.
Membrane porosity: Membrane porosity is determined as the ratio between the
volume of the pores and the total volume of the membrane. Evaporation surface
area increases with the increase in porosity level of the membrane, resulting in
higher permeate fluxes. Membrane porosity also affects the amount of heat loss by
conduction:
Qm = hm x ΔTm (2)
hm = ε x hmg + (1-ε) x hms (3)
where ε is the membrane porosity, hmg is the conductive heat transfer coefficient of
the gases entrapped in the membrane pores; hms is the conductive heat transfer
coefficient of the hydrophobic membrane material.
Conductive heat loss can be reduced by increasing porosity of the membrane, since
hmg is generally an order of magnitude smaller than hms. In general, the porosity of
the membranes used in MD operations lines in the range of 65%-85%.
Pore tortuosity: Tortuosity is the average length of the pores compared to
membrane thickness. The membrane pores do not go straight across the membrane
and the diffusing molecules must move along tortuous paths, leading a decrease in
MD flux. Therefore, permeate flux increases with the decrease in tortuosity. It must
be pointed out that this value is frequently used as a correction factor for prediction
of transmembrane flux due to the difficulties in measuring its real value for the
membranes used in MD. In general a value of 2 is frequently assumed for tortuosity
factor.
Membrane thickness: Permeate flux is inversely proportional to the membrane
thickness in MD. Therefore, membrane must be as thin as possible to achieve high
permeate flux. Thickness also plays an important role in the amount of conductive
heat loss though the membrane. In order to reduce heat resistances, it should be as
thick as possible leading to a conflict with the requirement of higher permeate flux.
Hence membrane thickness should be optimized in order to obtain optimum
permeate flux and heat efficiency. The optimum thickness for MD has been
estimated within the range of 30–60 μm.

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Pore size distribution: Pore size distribution affects uniformity of vapour permeation
mechanism. In general, uniform pore size is preferable rather than distributed pore
size.
Thermal conductivity: Thermal conductivity of the membrane should be small in
order to reduce the heat loss through the membrane from feed to the permeate
side. Conductive heat loss is inversely proportional to the membrane thickness.
However selection of a thicker membrane decreases both the flux and permeability.
One promising approach may be selection of a membrane with higher porosity since
thermal conductivity of polymer membrane is significantly higher than thermal
conductivity of water vapour in the membrane pores. The thermal conductivities of
polymers used in MD generally varies in the range of 0.15–0.45 [Wm-1
K-1
] depending
upon temperature and the degree of crystallinity.
2.4.1 Contact angle [2]
The contact angle is a common measurement of the hydrophobic or hydrophilic
behaviour of a material. It provides information about relative wettability of
membranes. The contact angle is determined as the angle between the surface of
the wetted solid and a line tangent to the curved surface of the drop at the point of
three-phase contact (Figure 4). The value of contact angle is greater than 90° when
there is low affinity between liquid and solid; in case of water, the material is
considered hydrophobic and is less than 90° in the case of high affinity. Wetting
occurs at 0°, when the liquid spreads onto the surface.
Fig.4, Schematic representation of contact angle, [2]
The wettability of a solid surface by a liquid decreases as the contact angle increases.
Table 4 lists the contact angle values for few different materials in water at ambient
temperature.

12
Table 4, Contact angle values of water on some materials at ambient temperature, [2]
2.4.2 Liquid entry pressure and wetting phenomena [2]
The hydrophobic nature of membranes used in membrane distillation prevents
penetration of the aqueous solutions into the pores unless a critical penetration
pressure is exceeded, as stated earlier. Liquid entry pressure (LEP) is the minimum
transmembrane hydrostatic pressure that must be applied before liquid solutions
penetrate into the membrane pores. LEP can be calculated using the Laplace-Young
equation,
ΔΡ = PF - PD - (2 x β x γL x cosθ)/ rm (4)
where PF and PD are the hydraulic pressure of the feed and distillate side, β is the
geometric pore coefficient (equals 1 for cylindrical pores), γL is the surface tension of
the liquid, θ is the contact angle and rm is the maximum pore size.
LEP depends on membrane characteristics and prevents wetting of the membrane
pores during MD experiments. LEP increases with a decrease in maximum pore size
at the surface and an increase at the hydrophobicity (i.e., large water contact angle)
of the membrane material. The presence of strong surfactants or organic solvents
can greatly reduce the liquid surface tension therefore causing membrane wetting.
Therefore, care must be taken to prevent contamination of process solutions with
detergents or other surfacting agents.
3. Transport mechanisms and polarization phenomena [2]
3.1 Theory of heat transfer
Heat transfer in the MD includes three main steps:
i. Heat transfer through the feed side boundary layer
ii. Heat transfer through the membrane
iii. Heat transfer through the permeate side boundary layer

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Heat transfer through the feed side boundary layer: Heat transfer from the feed
solution to the membrane surface across the boundary layer in the feed side of the
membrane module imposes a resistance to mass transfer since a large quantity of
heat must be supplied to the surface of the membrane to vaporize the liquid. The
temperature at the membrane surface is lower than the corresponding value at the
bulk phase. This affects negatively the driving force for mass transfer. This
phenomenon is called temperature polarization. Temperature polarization becomes
more significant at higher feed temperatures
The temperature polarization coefficient (TPC) is determined as the ratio of the
transmembrane temperature to the bulk temperature difference:
TPC = (Tfm - Tpm)/ (Tfb - Tpb) (5)
where Tfm, Tpm, Tfb and Tpb are membrane surface temperatures and fluid bulk
temperatures at the feed and permeate sides, respectively. A schematic diagram of
the temperature polarization in MD is shown in Figure 5.
Fig.5, Schematic diagram of temperature polarization in MD, Tfm, Tpm, Tfb and Tpb are membrane surface
temperatures and fluid bulk temperatures at the feed and permeate sides respectively, [2]
Heat transfer through the feed side boundary layer can be calculated using:
Qf = hf x (Tfb - Tfm) (6)
where hf is the heat transfer coefficient of the feed side boundary layer.
Heat transfer through the membrane: Heat transfer through the membrane appears
as a combination of latent heat of vaporization (QV ) and conductive heat transfer

14
across both the membrane matrix and the gas filled membrane pores (QC ). The
corresponding values can be estimated by following equations:
QV = J x ΔΗV (7)
QC = (km/δ) x (Tfm - Tpm) (8)
Therefore, the heat flux can be estimated by the following expression,
Qm = QV + QC (9)
Qm = (km/δ) x (Tfm - Tpm) + J x ΔΗV (10)
where km is the thermal conductivity of the membrane, δ is the membrane thickness,
J is the permeate water vapour flux and ΔHV is the latent heat of vaporization.
Various models have been proposed for estimation of km in Equation [10]. Two of the
most preferred ones are given below;
km = ε x kg + (1 - ε) x ks (11)
km = [ε/kg + (1 - ε)/ks]-1
(12)
Heat transfer through the permeate side boundary layer: Heat transfer from the
membrane surface to the bulk permeate side across the boundary layer is also
related with the temperature polarization phenomenon. The temperature of
membrane surface at the permeate side is higher than that of bulk permeate due to
the temperature polarization effect.
Heat transfer through the permeate side boundary layer is given as:
Qp = hp x (Tpm - Tpb) (13)
where hp is the heat transfer coefficient of the permeate side boundary layer.
Both feed and permeate side boundary layers are function of fluid properties and
operating conditions, as well as the hydrodynamic conditions. There are some
convenient approaches in the literature to reduce the temperature polarization
effects like mixing thoroughly, working at high flow rates or using turbulence
promoters.
3.2 Theory of mass transfer
As mentioned above, the mass transfer in MD is driven by the vapour pressure
gradient imposed between two sides of the membrane. Mass transfer in membrane
distillation consists of three consecutive steps:

15
i. Evaporation of water at the liquid/gas interface on the membrane surface of the
feed side
ii. Water vapour transfer through the membrane pores
iii. Condensation of water vapour at the gas/liquid interface on the membrane
surface of the permeate side
The mass flux (J) can be expressed as,
J = K x ΔΡ (14)
where K is the overall mass transfer coefficient which is the reciprocal of an overall
mass transfer resistance. This overall resistance is the sum of three individual
resistances:
K = [1/Kf + 1/Km + 1/Kp]-1
(15)
where Kf, Km and Kp are the mass transfer coefficients of feed layer, membrane and
permeate layer, respectively.
Mass transfer trough feed side boundary layer: In membrane distillation, only water
vapour transport is allowed due to the hydrophobic character of the membrane.
Therefore the concentration of solute(s) in feed solution becomes higher at the
liquid/gas interface than that at the bulk feed as mass transfer proceeds. This
phenomenon is called concentration polarization and results in reduction of the
transmembrane flux by depressing the driving force for water transport.
Concentration polarization coefficient (CPC) is determined as the ratio of the solute
concentration at the membrane surface (Cfm) to that at the bulk feed solution (Cfb):
CPC = Cfm/Cfb (16)
The concentration gradient between the liquid/gas interface and the bulk feed
results a diffusive transfer of solutes from the surface of the membrane to the bulk
solution. At steady state, the rate of convective solute transfer to the membrane
surface is balanced by diffusion of solute to the bulk feed.
The molar flux is expressed as follows,
J = ks x ln(Cfm/Cfb) (17)
where ks is the diffusive mass transfer coefficient through the boundary layer.
Several empirical correlation of dimensionless numbers, namely, Sherwood (Sh),
Reynolds (Re), Schmidt (Sc), Nusselt (Nu) and Prandtl (Pr) numbers can be used to
estimate the value of ks depending on the hydrodynamics of the system:
Sh = (k x L)/D, Re = (L x u x ρ)/μ, Sc = μ/ (ρ x D), Nu = (h x L)/k, Pr = (μ x Cp)/k (18)

16
where L: characteristic length, D: diffusion coefficient, ρ : density, μ: viscosity, u: feed
velocity, k: thermal conductivity, CP: specific heat, h: boundary layer heat transfer
coefficient.
In other membrane separation process such as microfiltration, ultrafiltration and
reverse osmosis, concentration polarization is usually considered a major cause for
flux decline. On the other hand, it is agreed upon that concentration polarization is
insignificant compared to temperature polarization in DCMD.
It is worth pointing out that in osmotic distillation process, concentration
polarization exists at each side of the membrane. During osmotic distillation, as mass
transfer proceeds, solute concentration increases at the membrane surface due to
evaporation of water vapour at the feed side. On the other hand, the solute
concentration decreases due to the condensation of water vapour on the permeate
side, giving rise to the difference in brine concentrations (Figure 6).
Fig.6, Schematic diagram of concentration polarization in MD. Cfm, Cpm, Cfb and Cpb are membrane surface and
bulk solute concentrations at the feed and permeate sides respectively, [2]
The existence of concentration polarization layers at each side of the membrane
results in the reduction of driving force for water vapour transport leading a
decrease in transmembrane flux.
Mass transfer through the membrane pores: The main mass transfer mechanisms
through the membrane in MD are Knudsen diffusion and molecular diffusion (Figure
7). Knudsen diffusion model is responsible for mass transfer through the membrane
pore if the mean free path of the water molecules is much greater than the pore size
of the membrane and hence, the molecules tend to collide more frequently with the
pore wall.

17
Fig.7, Mass transfer mechanism involved in water vapour transport through membrane pores of MD module,
[2]
In this case, the membrane diffusion coefficient is calculated using equation:
Κm = 1,064 x (r x ε)/(τ x δ) x (M/(R x T))0,5
(19)
where ε is the fractional void volume, δ is the membrane thickness, τ is the
tortuosity, M is the molecular weight of water, R is the gas constant and T is the
absolute temperature.
On the other hand, when the pore size is relatively large, the molecule–molecule
collisions are more frequent and molecular diffusion is responsible for mass transfer
through the membrane pores
Km = 1/Yln x (D x ε)/(Τ x δ) x (M/(R x T)) (20)
where Yln is the log mean of mole fraction of air and D is the diffusion coefficient.
Both models were successfully applied for predicting the mass transfer through the
membrane in DCMD systems.
4. Process parameters [2]
Feed concentration; Permeate flux decreases with an increase in feed concentration.
This phenomenon can be attributed to the reduction of the driving force due to
decrease of the vapour pressure of the feed solution and exponential increase of
viscosity of the feed with increasing concentration. The contribution of
concentration polarization effects is also known, nevertheless, this is very small in
comparison with temperature polarization effect. As it is well known, MD can handle
feed solutions at high concentrations without suffering the large drop in
permeability observed in other pressure-driven membrane processes and can be

18
preferentially employed whenever elevated permeate recovery factors or high
retentate concentrations are requested (i.e. concentration of fruit juices).
Feed temperature; Various investigations have been carried out on the effect of the
feed temperature on permeate flux in MD. In general, it is agreed upon that there is
an exponential increase of the MD flux with the increase of the feed temperature. As
the driving force for membrane distillation is the difference in vapour pressure
across the membrane, the increase in temperature increases the vapour pressure of
the feed solution, thus results an increase in the transmembrane vapour pressure
difference. It is worth quoting that working under high feed temperatures was
offered by various MD researches, since the internal evaporation efficiency (the ratio
of the heat that contributes to evaporation) and the total heat exchanged from the
feed to the permeate side is high. Nevertheless, the increase in quality losses and
formation of unfavorable compounds (i.e. hydroxymethyl furfural and furan) in fruit
juices due to high operation temperatures restricts the temperature levels.
Temperature polarization effect also increases with the increase in feed
temperature.
Feed flow rate & stirring; In MD, the increase in flow and/or stirring rate of feed
increases the permeate flux. The shearing forces generated at high flow rate and/or
stirring reduces the hydrodynamic boundary layer thickness and thus reduce
polarization effects. Therefore, the temperature and concentration at the liquid-
vapour interface becomes closer to the corresponding values at the bulk feed
solution. Onsekizoglu studied the effects of various operating parameters on
permeate flux and soluble solid content of apple juice during concentration through
osmotic distillation (OD) and membrane distillation (MD) processes. They observed
that the effect of feed flow rate on transmembrane flux was less than half of the
influence of temperature difference across the membrane. The effect of flow rate on
MD flux becomes more noticeable at higher temperatures especially associated with
higher temperature drop across the membrane. Consequently, higher productivity
can be achieved by operating under a turbulent flow regime. On the other hand, the
liquid entry pressure of feed solution (LEP) must be taken into account in order to
avoid membrane pore wetting when optimizing feed flow rate
Permeate temperature; The increase in permeate temperature results in lower MD
flux due to the decrease of the transmembrane vapour pressure difference as soon
as the feed temperature kept constant. It is generally agreed upon that the
temperature of cold water on the permeate side has smaller effect on the flux than
that of the feed solution for the same temperature difference. This is because the
vapour pressure increases exponentially with feed temperature.

19
Permeate flow rate; The increase in permeate flow and/or stirring rate reduces the
temperature polarization effect. Consequently, the temperature at the gas/liquid
interface approaches to the bulk temperature at the permeate side. This will tend to
increase driving force across the membrane; resulting an increase in MD flux. It is
important to note that as the permeate used in the MD is distilled water and in the
OD is hypertonic salt solution; the extent of the effect of flow rate is more prominent
in the latter configuration. This is because of the contribution of concentration
polarization effects on permeate side in OD.
4.1 Parameters to Reducing Temperature Polarization [4]
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 according to Equations (5), (6) and (10). The convective heat transfer coefficient
can be expressed as Equation (11):
αf = - k/ (Tfp - Tfm) x (dT/dy)boundary (21)
where k is thermal conductivity of the feed, and (dT/dy)boundary is the temperature
gradient in the thermal boundary layer of the feed. 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 9 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.
Fig.9, Spacer structure, [4]

20
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.
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.
5. Long-term performance; Membrane Fouling and Wetting
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 [4]:
• 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
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.
Membrane fouling & Cleaning procedures [2]; Membrane fouling refers to the loss
of membrane performance due to deposition of suspended or dissolved substances
on the membrane surface and/or within its pores. There are several types of fouling
in the membrane systems including inorganic fouling or scaling, particulate/colloidal
fouling, organic fouling and biological fouling (biofouling). Inorganic fouling or scaling
is caused by the accumulation of inorganic precipitates, such as calcium salts (CaCO3,
CaSO4), and magnesium carbonates on membrane surface or within pore structure.

21
Precipitates are formed when the concentration of these sparingly soluble salts
exceeds their saturation concentrations. Particulate/colloidal fouling is mainly
associated with accumulation of biologically inert particles and colloids on the
membrane surface. Organic fouling is related with the deposition or adsorption of
organic matters on the pores of the membrane surface. Microbial fouling however is
formed due to the formation of biofilms on membrane surfaces. Such films
(bacterial, algal, or fungal) grow and release biopolymers (polysaccharides, proteins,
and amino sugars) as a result of microbial activity.
Even though the general agreement is that the fouling phenomena is significantly
lower than those encountered in other pressure-driven membrane separation
processes, it is one of the major drawbacks in membrane distillation. The extensive
research on membrane fouling has revealed that the efficiency of MD installation
can be reduced by more than 50 percent after 50–100 h of process operation due to
the presence of fouling effects. In fact, all of the known types of fouling have been
determined to occur practically in MD operations.
Kullab & Martin (2011) pointed out that fouling and scaling may result pore clogging
in MD membranes, leading to a decrease in effective membrane area, and therefore
the permeate flux. Moreover, the flow channel area may be reduced resulting higher
temperature polarization due to the pressure drop across the membrane. The
increased deposition of the foulant species at the membrane surface would
eventually lead to an increase in the pressure drop to levels that the hydrostatic
pressure may exceed the LEP of the feed or permeate solution into the membrane
pores. Therefore the hydrophobic surface of membrane can be partially wetted due
to very small nature of the flow channels in MD modules (especially in hollow fiber
membrane modules).
Gryta (2005) presented the results of the over 3 years’ time research on the direct
contact membrane distillation applied for production of the demineralised water
using commercial capillary PP membranes. It was found that the membrane was
thermally stable, maintaining its morphology and its good separation characteristics
throughout the 3 years of DCMD operation. When using permeate of the RO system
as DCMD feed solution, membrane pore wetting was not observed; and the DCMD
flux was found to be similar to the initial permeate flux. However, precipitation of
CaCO3 on the membrane surface was observed when tap water was used directly as
a feed. A partial wetting of the membrane was found in this case resulting in a
decrease of the permeate flux from 700 to 550 L/m2
day. However, the formed
deposit was removed every 40–80 h by rinsing the module with a 2–5 wt% HCl
solution, permitting the recovery of the initial process efficiency. On the other hand,
authors reported that a multiple repetition of this operation resulted in a gradual
decline of the maximum flux of permeate.

22
Bubbling seems to be an obvious strategy to induce flow and improve shear stress at
the membrane surface to control polarization and fouling. Ding (2011) successfully
employed the intermittent gas bubbling method to reduce fouling layer formed in
concentrating TCM extract through DCMD. To limit membrane fouling or flux decline
during concentrating process, intermittent gas bubbling was introduced to the feed
side of membrane module. It was confirmed by experimental results that membrane
fouling was effectively controlled in the way of removing deposited foulants from
membrane surface by created two phase flow.
6. Engineering aspects; MD applications [2]
The MD process is currently applied mostly at the laboratory scale and the MD
applications are very appropriate for environmental, chemical, petrochemical, food,
pharmaceutical and biotechnology industries. Recently, some pilot plant applications
have been proposed for desalination and nuclear desalination but are still under
experimental tests and their use is not fully extended.
The major MD application has been in desalination for production of high purity
water. Near 100% rejection of non-volatile electrolytes (i.e., sodium chloride, NaCl;
potassium chloride, KCl; lithium bromide, LiBr; etc.) and non-electrolytes (i.e.,
glucose, sucrose, fructose, etc.) solutes present in aqueous solutions was achieved. A
quality water as low as 0.8 μS/cm electrical conductivity with 0.6 ppm TDS (total
dissolved solids) was produced. As the permeate product is very pure it is suitable
for use in medical and pharmaceutical sectors. In fact, in the case of a solution with
non-volatile components only water molecules flow through the membrane pores.
However, the obtained permeate fluxes were up to 1 kg/(m2
xh), which were lower
than the RO permeate fluxes (20-75 kg/(m2
xh)). Actually, due to MD membrane
module improvement, the MD production begins to be competitive to RO process in
the field of desalination with nearly total rejection factors, which cannot be
accomplished by RO at high permeate fluxes.
MD has been applied successfully to wastewater treatment at a laboratory scale,
either to produce a permeate less hazardous to the environment or to recover
valuable compounds. MD has been tested for the treatment of pharmaceutical
wastewater containing taurine, textile wastewater contaminated with dyes such as
methylene blue, aqueous solutions contaminated with boron, arsenic, heavy metals,
ammonia (NH3), coolant liquid (i.e., glycols), humic acid and acid solutions rich in
specific compounds, oil-water emulsions, olive oil mill wastewater for polyphenols
recovery and radioactive wastewater solutions. It was proved that DCMD is feasible
to process low and medium-level radioactive wastes giving high decontamination
factor in only one stage and can be applied for nuclear desalination. Recently, DCMD
was proposed for wastewater reclamation in space in a combined direct osmosis
system.

23
Due to the fact that MD can be conducted at relatively low feed temperatures, it was
successfully tested in many areas where high temperature applications lead to
degradation of the process fluids especially in food processing. It was demonstrated
that MD can be used for the concentration of milk, for the recovery of volatile aroma
compounds from black currant juice, for the concentration of must (i.e., the juice
obtained from grape pressing containing sugars and a wide variety of aroma
compounds) and for the concentration of many other types of juices including
orange juice, mandarin juice, apple juice, sugarcane juice, etc. It was concluded that
the utilization of either osmotic distillation (OD) and/or MD in the food industry for
concentration or separation is promising especially at high feed concentration
degrees.
MD also has potential applications in biotechnology. As an example, MD has been
used for the removal of toxic products from culture broths. The application of DCMD
unit connected to a laboratory bioreactor for the selective recovery of ethanol from
the culture medium has been reported. The experiments were run at a constant
temperature of 38o
C on anaerobic cultures of fragilis. MD was also applied for the
concentration of biological solutions such as bovine plasma and bovine blood. It was
demonstrated that MD was suitable for stable removal of solute free water from
blood with a haematocrit of 45%. DCMD was applied to the direct concentration of
protein (0.4% and 1% bovine serum albumin at pH 7.4) aqueous solutions at low
temperatures and found that fouling effects were practically absent, while the
limiting factor of the process was the temperature polarization.
It is known that azeotropic mixtures are impossible to be separated by simple
distillation. Thus, the application of MD for breaking azeotropic mixtures was
proposed and tested for the separation of hydrochloric acid/water, propionic
acid/water and formic acid/water azeotrope mixtures. It was demonstrated that MD
is of potential interest in breaking azeotropic mixtures. The effect of the inert gases,
helium, air and sulfur hexafluoride, in breaking the formic acid/water azeotropic
mixtures was studied. The selectivity was found to be larger and near unity when
using helium (around 0.96), followed by that in air (about 0.9) and then in sulfur
hexafluoride (0.85-0.86). The results were related with the different diffusivities of
the components in the inert gas.
MD has been proposed for the extraction of volatile organic compounds (VOCs) from
dilute aqueous solutions. Various types of dilute binary mixtures containing VOCs at
different concentrations were tested by different MD configurations and membrane
modules. Values of the selectivity different from those calculated on the basis of the
corresponding VLE data were found. Removal from water of alcohols such as
methanol, ethanol, isopropanol and n-butanol; halogenated VOCs such as
chloroform, trichloroethylene and tetrachloroethylene, benzene, acetone,

24
acetonitrile, ethylacetate, methylacetate and methyltertbutyl ether among others
were studied. The potential advantage of MD for ethanol recovery from
fermentation broth was also reported. It must be mentioned here that the addition
of salt such as magnesium chloride (MgCl2) during the treatment of aqueous alcohol
feed solutions was found to increase the alcohol selectivity significantly with only a
slight decrease in the total permeate flux. This was attributed to the reduction in
water vapour pressure leading to a decrease in the water mass transfer through the
membrane.
The concentration of aqueous solutions containing sodium hydroxide (NaOH) and
the strong mineral acid, sulfuric acid (H2SO4), at different pH values has been
investigated. Comparable MD permeate flux and electrical conductivity to those
obtained using sodium chloride (NaCl) aqueous solutions was noticed.MD separation
of aqueous solutions containing volatile solutes such as nitric acid (HNO3) and
hydrochloric acid (HCl) have been conducted and similar trends for both components
were found, different from that of the aqueous solutions containing non-volatile
solutes. Attempts were made for the concentration of hydrogen iodide (HI) and
sulphuric acid aqueous solutions in relation to hydrogen energy production from
water using DCMD and AGMD.
7. Advances on MD Processes and Modules for Water Purification [4]
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
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 10,
we see the simplest form of DCMD configuration which will desalinate a saline water
feed to a very high quality permeate.
Fig.10, Standard MD setup to desalinate water in direct contact mode, [4]

25
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 (22):
F = (1-t) x CP x (TF - TE)/ΔΗvap (22)
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%.
• Electrical energy constraints: The thermodynamics of the simple MD setup in turn
constrains the electrical consumption. Each pump in Figure 10 will consume
electrical energy per unit water permeated, Eelec,std (kWh/m3
), according to:
Eelec,std = PF/ (η x F) x 1/3600 (23)
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 10. 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

26
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:
• Fraunhofer ISE (AGMD);
• Memstill and Aquastill (AGMD);
• Scarab (AGMD);
• Memsys (vacuum enhanced multi effect AGMD).
7.3. Hybrid MD Systems
MD is a separation process that offers several unique features that conveniently
allow it to be integrated within other membrane operations. Most commonly, MD is
integrated into RO, nano-filtration (NF), and the more developmental forward
osmosis (FO).
7.3.1. MD Integration with RO or NF
One of the most logical technology partners for MD is RO or NF. There are two ways
in which they can be integrated. The first is by using the RO brine as feed to the MD,
or the NF or RO permeate as feed to the MD. These are represented in the flow
diagrams presented in Figure 11.

27
Fig.11, Simplified flow diagrams of hybrid RO/NF-MD systems. MD connected to RO concentrate (a) and to
RO/NF permeate (b), [4]
Using RO brine as a feed to MD (Figure 11a) has a great potential for MD utilization.
This directly addresses the upper concentration limit of RO at around 70,000 mg/L,
as MD is far less influenced by salt concentration. Typically, the need for an RO-MD
process to increase water recovery is for inland applications where disposal of the
brine is an issue. Testing of MD on RO groundwater concentrates revealed that the
concept is indeed viable, but suffers from practical issues such as scaling on MD
membranes. A similar result was found for an RO-MD trial on a solar powered direct
contact MD system in rural Victoria, Australia. Membrane scaling led to flux declines,
but flux was easily restored using an acid clean. Scaling was found to be effectively
managed by cleaning or the addition of anti-scalant. For the RO-MD process, the
individual RO recovery was 89%, and MDrecovery was 80%, giving a total water
recover of 98% for the combined system.
Integrating MD to treat RO or NF permeate (Figure 11b) is mostly concerned with
MD pretreatment. Scaling has been identified as a major issue for MD membranes
due to the capacity of scaling salts to “wet” the membrane (i.e., compromise the
membrane hydrophobicity leading to saline water leaking into permeate). To remove
scaling salts for water demineralisation applications (final water quality 1.5 to 2.5
μS/cm), Gryta tested tap water treated by NF prior to MD. While CaCO3 scaling
leading to flux decline was observed when treating the tap water directly by MD, HCl
cleaning removed scaling and restored full flux performance. To avoid this fouling
and cleaning issue, pretreatment using NF assisted the long term operation of MD,
but precipitation of a predominantly silica solids clogged the entrance of the module.
However, this was remediated by a simple filter at the module entrance.
7.3.2. MD Integration with FO
Forward osmosis (FO) is an emerging low pressure water treatment process that
relies on the natural osmotic force to transfer water through a semi-permeable
membrane from one solution to another. These solutions have differing dissolved
solid contents, which means that while the water has been taken from a non-potable
saline solution (e.g., seawater), it must be removed from the second solution (draw

28
solution) to become useable pure water. MD has been proposed for this second
removal step in novel space or protein concentration applications, schematically
represented in Figure 12. Although little explored, FO could recover water from a
brine with scaling salts such asgroundwater or seawater into a pure NaCl draw
solution. The draw solution is then reconcentrated byMD, and fresh water is
recovered.
Fig.12, Simplified flow diagram of FO-MD process for water desalination, [4]
8. Brackish water [4]
Two major reasons for needing to increase water availability all around the world,
are increasing water demands in urban centers with limited water resources, and the
over-pumping of fresh groundwater aquifers. As more communities diversify their
water sources, brackish groundwater and the use of membrane based processes has
gained significant traction. RO technology is mature and well understood; however,
its implementation for brackish water desalination is limited by two main drawbacks:
cost (capital and operating) and disposal of desalination concentrate. Brackish
groundwater is defined as “a source of water that exceeds the secondary drinking
water standard of 500 mg/L total dissolved solids (TDS) or the World Health
Organization (WHO) guidelines for drinking water quality of 1000 mg/L.” In the
United States, secondary standards are established for aesthetic purposes and as
such are not enforceable. Typical composition of brackish water as compared to sea
water and other impaired waters is presented in Table 5.
Table 5, Comparison of typical composition of produced water and other impaired waters for potential MD
applications, [4]

29
The salinity of the groundwater results from a dissolution process where chloride
(Cl), sulfate (SO4), sodium (Na), and calcium (Ca) are the dominant ions. The brackish
groundwater is blended with water from distillation plants to make the water
suitable for drinking. Treated water is also used to cover agricultural and domestic
needs. The number of inland desalination plants is growing considerably in Europe
and other parts of the world. As compared to most desalination plants that return
the salt concentrate to the ocean, inland plants must find other alternatives for
disposal and reduction of concentrate. Membrane distillation can be a
complementary technology to treat the brine waste generated by RO. As an
example, Macedonio and Drioli demonstrated that combining MD with RO operation
using a process intensification approach can increase the RO recovery factor and
extend the life of the RO membrane. In this approach, one portion of the RO
permeate is treated in the MD system instead of passing all the first stage RO
permeate through a second RO stage.
8.1 RO/ED/EDR Concentrate
Reverse osmosis of brackish groundwater (BWRO) has found increasing application
in semi-arid and arid countries to treat brackish groundwater for drinking, industrial,
or irrigation purposes. In the Middle East and the United States, RO treatment plants
have been implemented and are in operation. RO plants are also in operation in
Europe and Australia. The last one has six seawater RO plants for its major cities and
one to support mining operations. Additionally, Australia has a vapor compression
system to support mining operations. The Tampa Bay Seawater Desalination Facility
in Tampa, United States, which is the only large-scale facility in the country using a
coastal surface water source, operates using reverse osmosis. The largest seawater
RO plant in the world, located in Ashkelon, Israel, has a design capacity of 326 MLD.
The Sorek desalination plant in Israel to be completed in 2013 will have a design
capacity of 410 MLD.
One of the major concerns for BWRO is the disposal of the RO concentrate, arising
from the presence of anti-scalants, pre-treatment chemicals, and remoteness from
the sea or another economically viable concentrate disposal options. The volume of
concentrate produced depends on factors such as source water quality (e.g., salinity
level) and technology utilized. Table 6 presents examples of concentration of feed
water and corresponding RO concentrate for various brackish waters in the State of
Texas, United States and The United Arab Emirates.

30
Table 6, Examples of main composition in feed water and concentrate in desalination facilities (mg/L), [4]
In the United States the main concentrate disposal method is deep-well injection. In
Australia, most of the facilities dispose their concentrate via ocean outfall, although
smaller inland plants discharge to the sewer or evaporating basins, or use ground
infiltration. The presence of salt, metals, and silica at or above super-saturation due
to the addition of antiscalant and dispersants during the RO process may be a major
concern for disposal of desalination concentrate in deep well injection, since
eventually unwanted precipitates may form. A study conducted by Macedonio and
Drioli reported that combining RO with MD allowed total boron and arsenic rejection
from salty water without the need for addition of oxidizing agents, resulting in less
environmental impact.
8.2 Concentrate Management Cost for MD
The cost of concentrate disposal can be significant. This is particularly valid for
brackish water RO plants that use groundwater as the source water and have to
dispose of the concentrate in either a different ground formation to ensure that it
would not migrate and potentially contaminate a source of drinking water or in lined
evaporation ponds. For these plants surface water outfall is not a disposal option.
Even with advances in membrane production costs to have less expensive
membranes, longer membrane life, and energy recovery improvements, the cost of
concentrate management represents an increasing percentage of the total water
treatment plant cost. Table 7 shows that the cost for pretreatment and RO
treatment of produced water increases as the concentration of the concentrate
increases.

31
Table 7, Power cost of desalination in oil field brine, [4]
Based on current concentrate disposal limitations, reclamation of effluents for
irrigation and indirect potable water uses is rapidly developing as an alternative to
seawater desalination. Cote compared the total life cycle costs for treating water
from secondary effluent using RO and seawater desalination, and found that they
are $0.28/m3
and $0.62/m3
, respectively.
An alternative disposal option for concentrate management is to treat the
concentrate to recover potential economic products and have zero liquid discharge.
Zero liquid discharge presents an opportunity for MD in that, being a thermal
process, it can concentrate saline water to the precipitation of salts with minimal
drop in flux performance. By analyzing the cost of disposal in brine lagoons in Table
22, it is observed that technologies to reduce the disposal volume have good
economic incentives. Capital cost included reductions estimated using RO or
combined RO and MD desalination. Data in Table 8 assumes pond price at $1
million/ha (including pond lining), evaporation rate of 1.0 m/year, and 75% RO water
recovery. Even with RO taking to 70,000 mg/L at its limit, further concentrating the
brine beyond this limit is worth a savings of $17 million for a 5 MLD plant.
Table 8, Example of brine disposal pond capital for feed water stream of 5 ML/day, [4]

32
Besides this financial incentive, environmental incentives likewise drive zero liquid
discharge as brine disposal to lagoons may not be an acceptable practice due to risk
of uncontrolled saline water release. MD is a potential treatment candidate for
coupling with RO or ED to increase water recovery and to reduce the amount of
concentrate requiring disposal to get closer to zero liquid discharge. Thermal
desalination by MED would compete with MD in this space. However, when low cost
thermal energy is available, MD can be an economical alternative to the established
thermal processes in zero liquid discharge applications. MD is a technology that can
be coupled with RO and/or others to reduce waste streams (i.e., RO concentrate). It
could also be used for small-scale applications in which the water quality is not
suitable for RO based processes. MD can also be co-located with power generation
facilities or industrial facilities to take advantage of the waste heat to produce high
quality water.
Additionally, MD can use a variety of membranes, which clearly presents a variation
on the cost of the treated water. While polypropylene (PP), polyvinylidenefluoride
(PVDF) and polytetrafluoroethylene (PTFE) are the most widely used membrane
materials, Their prices vary not only on the original material prices, but also in their
design and performance. A survey of the materials has been conducted, and PTFE
was found to offer best performance due to its highly hydrophobic character. Also,
the support layer was found to greatly influence performance, with improved MD
performance obtained by membranes supported on woven scrim materials. Low
cost, high quality membranes emerging from China at present have a price less than
$10 per m2
, playing a key role in making MD affordable.
9. Investigation of high recovery of concentrated RO brine using MD [3]
Eastern Municipal Water District (EMWD) in Southern California has implemented
the Perris Basin Desalination Program to reduce its dependence on a costly and
potentially limited supply of imported water. In order to utilize high-TDS
groundwater from its basins, EMWD is operating two reverse osmosis (RO)
desalination facilities and designing a third. The groundwater is blended with RO
product water from the facilities to achieve product water with less than 500mg/L
TDS in the distribution system. The RO brine stream is discharged into the 22-mile-
long Temescal Valley Regional Interceptor, which is a non-reclaimable waste pipeline
that connects EMWD to the Santa Ana Regional Interceptor (SARI). The brine is then
transported by the SARI to Orange County Sanitation District (OCSD) for treatment
and discharge. Operation of all three RO facilities will ultimately produce brine
quantities in excess of EMWD’s capacity in the SARI system, and additional capacity
is not available. Furthermore, the cost of treatment and disposal by OCSD is
expected to increase. Therefore, like many other inland water utilities, EMWD must
improve water recovery.

33
Additional brine treatment to approach zero liquid discharge (ZLD) would not only
enable EMWD to produce more water, but also to reduce their reliance on existing
brine disposal methods. In 2005, the California Department of Water Resources and
the United States Bureau of Reclamation sponsored a study at EMWD with the
objectives of increasing water recovery and decreasing brine volume.
Abbreviations: CF, concentration factor; CP, concentration polarization; CTA, cellulose
triacetate; ED, electrodialysis; EDR, electrodialysis reversal;
Fig.13, Schematic drawing of the bench-scale VEDCMD system, [3]
9.1. Brine chemical analysis
Two RO brines were generated during the investigation. The first brine was
concentrate from the primary RO process. Water recovery in the primary RO system
was limited to 70% to avoid precipitation of sparingly soluble salts on the
membranes. As part of the effort to achieve higher recovery, the primary RO brine
was softened and further treated in an electrodialysis reversal (EDR) system or a
secondary RO system, thus generating the second brine.
The compositions of brines A and B are summarized in Table 9. The TDS
concentration of brine B is approximately 2.5 times greater than that of brine A.
Considering individual ions, sodium and chloride ion concentrations in brine B are
higher than in brine A, while calcium, sulfate, and silica concentrations in brine B are
lower than in brine A. These differences are due to the softening treatment of brine
A before it was fed to the secondary RO to generate brine B.

34
Table 9, Water qualities of brines A and B, [3]
The softening treatment was used to reduce calcium, sulfate, and silica
concentrations in brine A to prevent scale formation on the RO membranes during
secondary RO. The reductions in calcium, sulfate, and silica concentrations and the
concentration process in the secondary RO system resulted in the increase in relative
concentrations of sodium and chloride in brine B.
Salts that exceed their saturation and precipitate out of solution do not affect
osmotic pressure, but will cause scaling of the membrane. A chemical simulation
program was used to determine the saturation level of the potentially scaling
minerals. In brine A, SiO2 and CaSO4 were found to be at 99 and 89% saturation,
respectively; in brine B, SiO2 and CaSO4 were found to be at 57 and 50% saturation,
respectively. Thus, it was anticipated that membrane scaling would occur earlier
during experiments with brine A. Furthermore, because the solubility of CaSO4 is
inversely proportional to temperature (decreasing solubility with increasing
temperature); a higher percent saturation would be expected for higher brine
temperatures. However, the decrease in solubility (approximately 200 mg/L from 40
to 60 ◦C) results in a 0% change in percent saturation; thus, inverse temperature
effects are essentially negligible for the feed temperatures (40 and 60 ◦C) in the
current VEDCMD investigation.
9.2. Vacuum enhanced direct contact MD
9.2.1. Water flux and recovery
Water flux as a function of concentration factor (CF) is illustrated in Fig. 3a and b for
VEDCMD of brines A and B, respectively;

35
Fig.14, Water flux as a function of CF and batch recovery for VEDCMD of (a) brine A and (b) brine B using the
PTFE membrane. Initial feed concentration in brine A was 7500 mg/L TDS and in brine B 17.500 mg/L TDS, [3]
batch recovery is also shown on the top x-axes. CF is the ratio between the
concentration of the feed solution at any time and the initial feed concentration.
Batch recovery is the cumulative volume of permeate collected during an
experiment until a point in time normalized to the initial feed volume. CF is related
to batch recovery by CF = 1/(1−R); thus, CF and R are not linearly related.
The PTFE membrane was the only membrane initially tested because it was known
to have higher permeability, and therefore higher water flux than the PP membrane.
The experiments were terminated when water flux reached approximately 5 L/(m2
h). During all experiments, salt rejection was greater than 99.9%.
It is apparent from the results in Fig. 3a and b that flux decline is substantial for
almost all experimental conditions. When comparing the flux declines in these
experiments with those from a previous investigation (black line in Fig. 3a), it can be
seen that water flux decreases much more rapidly in the current investigation. In the
previous investigation, water flux decreased only slightly with increasing feed TDS
concentration because the salts studied were NaCl and sea salt—neither of which
contains ions that are likely to exceed their solubility and form scale on the
membrane in the range of feed concentrations tested.
Also in Fig. 3a and b, initial water fluxes were substantially greater in experiments
conducted with a temperature difference of 40 ◦C than those conducted with a
temperature difference of 20 ◦C. A higher temperature difference results in a higher
vapor pressure difference across the membrane and a stronger driving force for
water evaporation. Initial water fluxes were even higher when the permeate

36
pressure was lowered from 660mmHg (abs) to 360mmHg (abs). A lower permeate
pressure results in a higher partial vapor pressure difference and an increased
driving force.
In experiments conducted with brine A (Fig. 14a), a relatively constant water flux was
observed up to a CF of approximately 1.75, at which point a rapid flux decline was
observed for all experiments. These flux declines were preceded by observed
changes in feed water clarity at CFs of approximately 1.5; this was likely evidence
that silica and calcium sulfate had exceeded their saturation and were precipitating
out of solution. Thus, the rapid flux declines were likely caused by precipitated solids
(mineral scale) on the membrane surface that blocked the transport of water vapor
through the pores. SEM micrographs of the membrane surfaces after experiments
revealed thick scale layers on the membrane (Fig. 15). Calcium sulfate crystals and
some silica aggregates are shown in the inset micrograph,
Fig.15, An SEM micrograph of a cross-section of the scaled PTFE membrane showing the membrane support
layer, active layer and calcium layer (CaSO4) scale after VEDCMD of brine A, [3]
The fact that the onset of flux decline in Fig. 14a and b is earliest and most rapid in
the experiments that started with the highest initial flux (ΔT=40◦C, Pp = 360mmHg
(abs)); and latest and most gradual in the experiments that started with the lowest
initial flux (ΔT=20◦C, Pp = 660mmHg (abs)) can also be explained by the exponential
relationship between water flux and concentration polarization (CP) at the feed–
membrane interface. ‘CP increases exponentially with increasing water flux
according to the classical film model,
CPmodulus = Cm/Cb = (1-Ro) + Ro x eJ/k
(24)
where Cm is concentration at the membrane, Cb is concentration in the bulk feed
solution, R0 is the observed salt rejection, J is the permeate flux, and k is the solute
mass transfer coefficient on the feed side. Therefore, at higher water fluxes, the

37
increased solute concentration near the membrane surface would cause SiO2 and
CaSO4 to exceed their solubility and form scale on the membrane. Along the same
lines, the highest batch recovery was achieved in the experiment with the lowest
initial flux (ΔT=20◦C and Pp = 660mmHg (abs)).
When comparing results from experiments conducted with brine A (Fig. 14a) to
those conducted with brine B (Fig. 14b), it is apparent that TDS concentration has
minimal effect on initial water flux—a substantial advantage over pressure-driven
membrane desalination processes. Also, when comparing results of brine A and
brine B, similar trends in flux with time were observed. Much higher CFs (or batch
recoveries) were achieved for brine B (Fig. 14b) than for brine A (Fig. 14a). Higher
batch recoveries were anticipated for Brine B based on the water quality data (Table
9) and percent saturation values that were lower in brine B due to the softening
process. Also, brine B contained residual scale inhibitor, which was used to inhibit
the formation of CaSO4 and SiO2 during the secondary RO treatment.
9.2.2. VEDCMD membrane cleaning
One of the objectives of the study was to investigate the ease by which the scale
layers could be removed from the membrane surface. The membranes were
chemically cleaned with Na2EDTA solution after their water flux dropped below 5
L/(m2
h). Brine A was used as the feed in these experiments because it scaled the
membrane more rapidly than brine B (Fig. 14). Also, in order to expedite scale
formation, experiments were conducted with a temperature difference of 40 ◦C
instead of 20 ◦C.
Water flux and batch recovery before and after membrane cleaning are shown in Fig.
16a and b for the PTFE and PP membranes, respectively.
Fig.16, Water flux as a function of time in VEDCMD cleaning experiments with (a) the flat-sheet PTFE
membrane and (b) the flat sheet PP membrane. Brine A feed solution with Tf = 60
o
C, Tp = 20
o
C and Pp = 660
mmHg, [3]

38
The PP membrane was included in the cleaning experiments to compare its fouling
tendency and chemical resistance with the PTFE membrane. The performance of the
PTFE membrane (Fig. 16a) was different before and after cleaning. The initial water
flux after cleaning was the same as the initial flux before cleaning, except that after
cleaning there was an immediate onset of flux decline. This suggests that the
majority of scale was removed from the membrane following cleaning, thus
restoring water flux to its initial level; however, the residual scale that did remain on
the membrane most likely provided sites for crystallization, leading to more rapid
scale formation and earlier onset of flux decline after cleaning. The performance of
the PP membrane (Fig. 16b) was similar before and after cleaning. This implies that
scale deposit on the PP membranes is less strongly adhered to the membrane and
can be removed using a simple cleaning method.
Both the PTFE and PP membranes are characterized as having high chemical
resistance. To ensure that exposure to the EDTA cleaning solution did not damage
the membranes, their rejection was monitored throughout the experiments; both
membranes maintained greater than 99.9% salt rejection before and after cleaning.
This suggests that the PTFE and PP membranes are indeed chemically resistant to
the EDTA over short terms.
9.2.3. VEDCMD with scale inhibitor
The effect of dosing brine A with CaSO4 scale inhibitor (Pretreat Plus 0400) was
investigated in a separate set of VEDCMD experiments. The experiments were
conducted under low flux conditions (ΔT=20◦C, 660mmHg(abs)) using the PTFE
membrane. Water flux as a function of CF and batch recovery is illustrated in Fig. 17.
For reference, a VEDCMD experiment was performed without addition of scale
inhibitor (solid black line).
Fig.17, Water flux in VEDCMD of brine A as a function of CF for different doses of calcium sulfate scale
inhibitor. Brine A feed solution with Tf = 60
o
C, Tp = 20
o
C, Pp = 660 mmHg and initial feed concentration of 7500
mg/L, [3]

39
A very rapid flux decline followed by partial recovery was observed during all of the
experiments with the scale inhibitor. Although the flux recovery is not fully
understood, the unusual flux behavior is likely due to the formation of amorphous
silica. Amorphous silica and silicates precipitate in a series of steps generating soft
and then hard gels. In the current investigation, it is likely that the rapid flux decline
was due to the formation of soft silica gels that form on the membrane surface.
Further reaction of the silica resulted in hard gels that were scoured off of the
membrane surface; this resulted in the flux recoveries observed in Fig. 6. It can also
be seen in Fig. 17 that the highest batch recovery occurred with a scale inhibitor
dose of approximately 4 ppm, yet, further optimization and better understanding of
the chemical and physical phenomena are needed.
9.2.4. VEDCMD water recovery
Approximately 62% batch recovery was achieved for VEDCMD of brine A using the
PTFE membrane at the lowest flux conditions (ΔT=20◦C, 660mmHg (abs)) (square
symbols in Fig. 14a). The use of scale inhibitor brought the water recovery to
approximately 78% (Fig. 17). In comparison, greater than 80% batch recovery was
achieved for VEDCMD of brine B using the PTFE membrane at the lowest flux
conditions (ΔT=20◦C, 660mmHg (abs)) (square symbols in Fig. 14b).
In order to determine the total water recovery, the recovery from EMWD’s RO
processes and the batch recovery from the VEDCMD process were both considered.
The total recovery was calculated using:
Rtot = RRO + (1 - RRO) x RVEDCMD (25)
where Rtot is the total water recovery, RRO is the water recovered from EMWD’s RO
processes, and RVEDCMD is the batch recovery from the current study. For brine A, RRO
was 70%andRVEDCMD was 62%. For brine B, RRO was 89% and RVEDCMD was 80%. Thus,
when combining the recoveries of the RO processes and the VEDCMD process, the
total recovery was greater than 89% for brine A and greater than 98% for brine B.
9.3 Comparing VEDCMD and FO for brine treatment
A comparison between VEDCMD and FO of brine A is shown in Fig. 18a.

40
Fig. 18, Comparison of high and low temperature VEDCMD and FO for (a) brine A and (b) brine B. Initial feed
concentration in brine A was 7500 mg/L TDS and in brine B 17.500 mg/L TDS, [3]
High temperature VEDCMD (40 ◦C) had a substantially substantially higher initial flux
than low temperature VEDCMD (20 ◦C) or FO. For both VEDCMD experiments, rapid
flux decline is observed due to scale depositing on the membrane surface. Flux
decline is more gradual during the FO experiment. This suggests that scaling was not
as severe in FO as it was in VEDCMD. The highest water recovery was achieved using
FO; with a batch recovery of 87%, it substantially outperformed the VEDCMD
processes. A comparison between VEDCMD and FO of brine B is shown in Fig. 18b.
Again, high temperature VEDCMD had a substantially higher initial flux than low
temperature VEDCMD or FO. For high temperature VEDCMD, a steep flux decline
was observed; for low temperature VEDCMD and FO, relatively gradual flux declines
were observed. The water recovery from brine B was low using FO because the high
ion concentration in the brine substantially increased the feed osmotic pressure,
which in turn reduced the osmotic pressure difference, and thus the driving force
across the FO membrane. The highest water recovery was achieved using low
temperature VEDCMD; with a batch recovery of 79%, it substantially outperformed
both high temperature VEDCMD and FO.
9.4 Conclusions
In this study, it was found that FO outperformed low- and high-temperature
VEDCMD when treating a feed with high scaling propensity but low TDS
concentration (i.e., brine A); and low temperature VEDCMD outperformed high-
temperature VEDCMD and FO when treating a feed with lower scaling propensity
but high TDS concentration (i.e., brine B). High temperature VEDCMD results in
higher initial water flux, but also greater flux decline. In FO, the high osmotic
pressure of the feed solution coupled with the scaling environment may limit the
utilization of the process for desalination of highly saline source waters; however,
new draw solutions and methods of re-concentration could alleviate the low
performance observed when treating feeds with high osmotic pressure. In all
experiments, scale formed on the active surface of the membranes and adversely

41
affected batch recovery, but cleaning methods were effective at removing scale from
both the MD and FO membrane surfaces. It was also found that by dosing the
feedwater with an appropriate scale inhibitor, a substantial improvement in batch
recovery for both VEDCMD and FO could be achieved. When considering the total
water recovery (the recovery from the RO processes combined with the batch
recovery from the VEDCMD or FO process), greater than 96 and 98% total recoveries
were achieved for the two different brine streams.
9.5 Recommendations
This work demonstrates the potential benefits of MD processes for the minimisation
of brine wastes. To further develop the technology the following areas require
investigation:
• Optimisation of the membrane module and spacer design to maximise water flux
for a given input temperature difference. This was identified as a likely source of
lower than expected flux.
• Identify the main foulants observed on the membrane and investigate strategies to
combat fouling. Removal of precipitate using in-line filtration may significantly
improve long term performance. Simple pH control also may show fouling
improvements.
• Determine what influence antiscalant chemicals used in conventional RO systems
have on the operation of MD membranes with respect to scaling, fouling and flux.
• Scale up to semi-pilot dimensions to investigate limitations of larger membrane
areas and include heat recovery equipment to determine economic feasibility of
potential commercial installations
• Undertake research into the extent of liquid water protrusion into membrane
pores to predictoptimal membrane thickness.
• Identify obstacles to further scale-up to a demonstration size MD plant

42
10. Membrane distillation as a means for reverse osmosis concentrate
volume minimization [10,11]
Membrane distillation (MD) is a non-isothermal evaporative technology that uses a
hydrophobic microporous membrane being the driving force the vapor pressure
difference between both membrane sides. MD can be applied for the treatment of
saline solutions with high concentrations.
MD is commercially available and produces very high-quality distillate [10];
Advantages
- salt rejections of 99–100% are achievable in most circumstances
- the feedwater does not require the extensive pretreatment that is typically vital for
pressure-based membrane processes, which makes it technically feasible for treating
large amounts of water in seawater desalination plants
- energy requirements are high relative to energy use of RO, but less than traditional
evaporation and crystallization systems
- as the driving force for MD is temperature difference, very low feed temperatures
can produce reasonably high rates of product water and may be more practical
considering the nature of some water impurities (e.g. scaling issues at high
temperature)
- Low feed temperatures also allow the use of low-grade heat such as industrial
waste heat, solar or desalination waste heat, so that MD can be easily coupled with
solar ponds
Disadvantages
- MD could have problems related to scaling & fouling on the membranes.
[11] To increase the flux of MD systems, several polymers have been studied in the
past. Polypropylene (PP), polytetrafluoroethylene (PTFE), polyacrylonitrile (PAN) and
polyvinylidenefluoride (PVDF) have been commonly employed in the preparation of
membranes for MD applications. Compared to dual-layer hydrophobic– hydrophobic
PVDF and dual-layer hydrophobic–hydrophilic PVDF/PAN membranes, the single-
layer PVDF membrane exhibited lower reduction in membrane permeability due to
membrane morphology and pore size, which played a more important role than
membrane wall thickness. In one study, Dumιe introduced carbon nanotubes within
the MD polymer matrix to increase permeability. Incorporation of carbon nanotubes
resulted in higher contact angles (113o
), higher porosity (90%) and lower thermal
conductivity when compared to polymeric MD.

43
Apart from improving mass transfer, enhancing flux in MD systems by controlling
scaling and fouling is also important. Integrated systems, where MD is used to treat
the concentrate from NF/RO membranes, have been studied. Ji combined MD with
crystallization (MDC) for treatment of brines discharged from a seawater RO system
and obtained an initial flux of 1.44 L m-2
h-1
and a feed water recovery up to 90%
with the simultaneous production of NaCl crystals. The initial flux decreased by 13%
due to natural organic matter (NOM) fouling of the MD membrane. The presence of
NOM in the RO concentrate also affected NaCl crystallization kinetics in terms of
reduced magma density, nucleation and growth rates. Mericq utilized vacuum
membrane distillation (VMD) to treat RO concentrate during seawater desalination
and achieved an overall recovery of 89%.
To prevent crystallization near the membrane surface, Creusen utilized osmotic
distillation. In this approach, a draw solution (such as CaCl2) is introduced in the
distillate side causing a decrease in vapor pressure and an inverse temperature
profile(reduced temperature polarization). Thus, the heat of evaporation on the feed
side is provided by the distillate side and a temperature drop occurs at the feed side
preventing crystallization. In another approach, Nghiem and Cath altered the
induction time to reduce scaling by CaSO4. In this approach, periodic flushing of the
membrane with permeate reset the induction time of CaSO4 and resulted in effective
scale control. At low system recoveries, the permeate flux was constant even at
super saturation conditions.
10.1 Comparison of emerging technologies for concentrate treatment
Comparison of emerging technologies for concentrate treatment is presented in
Table 10. The FO process achieves similar treated water quality when compared to
other membrane technologies for concentrate treatment but limited full-scale
applications exist. Although the energy consumption of the membrane process in FO
requires lower energy consumption, recovery of the draw solution results in the
overall specific energy consumption similar to other membrane processes. A
particular advantage with the FO process is the higher limits on TDS of concentrate
that requires treatment. Unlike the RO process, the FO process can be utilized for
treating concentrate streams with a TDS of up to 100,000 mg/L. The MD process has
similar advantages for treatment of concentrate with high TDS levels. A particular
advantage of the FO and MD process is in the presence of a waste heat source to
heat the feed water to MD or regenerate the draw solution in FO. All the emerging
technologies addressed in this review do not require applied pressure to treat the
concentrate but the technologies are still under developmental stages and have
been applied only at pilot or demonstration scale. Cost estimates for emerging
technologies are not available as these technologies have been evaluated only at the
pilot-scale.

44
Table 10, Comparison of emerging technologies for RO concentrate treatment, [11]
a. Vacuum membrane distillation (VMD) is a variant of MD, in which low pressure or
vacuum is applied on the permeate side of the membrane module, for example by
means of vacuum pump(s). The applied permeate pressure must be lower than the
saturation pressure of volatile molecules to be separated from the feed solution and
condensation takes place outside the membrane module at temperatures lower
than the ambient temperature.
Fig.19, Schematic of seawater desalination by RO and vacuum membrane distillation (VMD) integrated
process,[11]
Mericq applied VMD configuration for the treatment of synthetic RO brines
containing only the mineral part of seawater with total salt concentrations up to
300 g/L. High permeate fluxes were obtained even for the highest salt
concentrations. However, the permeate flux was limited at high salt concentrations
by scaling, mainly due to calcium precipitation.
Despite this inconvenience, scaling had only a partial impact on the permeate flux
(i.e. 24% decrease for 43 L/(m2
·h) for the permeate with the highest salt
concentration. Calcium carbonate (CaCO3) and calcium sulfate (CaSO4) precipitated

45
first due to their low solubility and formed mixed crystal deposits on the membrane
surface. These phenomena only occurred on the membrane surface and did not
totally cover the membrane pores. The crystals were easily removed simply by
washing the membrane with water. Simulations were performed to study the yield
of the process with 40,000 m3
/day of 38.9 g/L seawater, achieving a recovery of
40% for VMD itself and up to 89% for overall recovery by coupling RO and VMD.
Results also showed that concentrate quantity can be reduced by a factor of 5.5,
making it possible to double overall water production.
b. Membrane distillation crystallization (MDC)
Ji investigated the performance of membrane distillation crystallization (MDC) at
bench-scale in terms of water recovery and NaCl crystallization kinetics. The
extensive contact area provided by hollow fiber membranes made it possible to
achieve reliable permeate fluxes at moderate temperatures (40–50 °C) with energy
consumption ranging from 15 to 20 kWh/m3
,which is lower than that of
conventional evaporative systems for NaCl crystallization having a specific energy
consumption of 30 kWh/m3
. Experimental tests carried out on artificial RO
concentrates resulted in 21 kg/m3
production of NaCl crystals and the final water
recovery factor increased up to 90%. Analogous investigations carried out on RO
brines from natural seawater were affected by the presence of dissolved organic
matter, showing 20% reduction in the amount of salt crystallized and 8% decrease
of the permeate flux. Therefore, adequate pretreatment before the RO stage is
needed to reduce the negative effect of dissolved organic matter on the MDC
performance. This study confirms the ability of MDC to concentrate RO brines. In
principle, the industrial scale-up of the MDC process involving large volumes of
brines do not show any technical complexity.
c. Vacuum-enhanced direct contact membrane distillation (VEDCMD)
Martinetti studied vacuum-enhanced direct contact membrane distillation
(VEDCMD) to increase water recovery during desalination of brackish water (Fig. 20).

46
Fig.20, Schematic drawing of a vacuum-enhanced direct contact membrane distillation (VEDCMD),[11]
VEDCMD differs from VMD in its additional direct contact system, in which warmer
feedwater is in contact with the active side of the membrane and a cooler water
stream is in direct contact with the support side. In their tests, two RO brine streams
were used as feed of the VEDCMD system, with total dissolved solid
concentrations ranging between 7500 and 17,500 mg/L. A recovery factor up to
81% was achieved. However, recovery factors were always limited by the
precipitation of inorganic salts on the membrane surface. Martinetti showed also
that cleaning techniques were able to remove the scaling layer from the membrane
surfaced restoring the water permeate flux to almost its initial level. The authors also
claimed that the addition of scale inhibitors during the process was effective in
maintaining high water permeate flux during an extended VEDCMD operating time.
d. Salt-gradient solar pond
A salt-gradient solar pond is a body of saline water in which the salt concentration
increases with depth, from a very low value at the surface to near saturation at the
bottom. The density gradient inhibits free convection, and the result is that solar
radiation is trapped in the lower region. Lu provided heat to MD systems with a
coupled salt-gradient solar pond. The MD unit was successfully operated at a first-
stage vapor temperature range of 60–75 °C, and at a very high concentration ratio
with the reject brine near saturation. The temperature level has a significant effect
on both production rate and performance ratio. The production rate increases, but
the performance ratio decreases with both increased temperature and increased
temperature differences between the first and fourth stages. The membrane
distillation unit produces high-quality distillate of about 2–3 mg TDS/L. Quiblawey
did an overview of solar thermal desalination technologies focusing on those
technologies appropriate for use in remote villages and concluded that solar energy
coupled with desalination offers a promising prospect for covering the
fundamental needs of power and water in remote regions, where connection to

47
the public power grid is either not cost-effective or not feasible, and where water
scarcity is severe.
11. PRO concentrate treatment with DCMD [6]
Three kinds of PRO concentrate (direct PRO concentrate, silica bearing PRO
concentrate, and iron and manganese bearing PRO concentrate) were used in the
experiments. The RO was fed with the tap water with the recovery of 50%. After
DCMD process, the whole water recovery can be significantly enhanced to 98.8%.
Membrane clogging caused by the formed deposit (CaCO3, CaSO4, and silicate) was
the main reason of the membrane efficiency reduction. It was found that CaCO3
formed at first during the process, and can be alleviated even eliminated by
acidification. When the PRO concentrate was concentrated a high level, CaSO4
formed and caused a sharp decline of module efficiency. During the DCMD process
of silica bearing PRO concentrate, silica may co-precipitate with soluble metals to
form silicate at alkaline solutions while colloid silica may form at acidic solutions.
Fig.21, Membrane distillation setup;(1) feed reservoir;(2) membrane module; (3) permeate reservoir; (4) water
bath; (5) cooling oil; (6) pump; (7) thermometer and (8) conductivity monitor, [6]
The membrane module was made by a polyester tube and two UPVC T-tubes. The
module was equipped with hydrophobic hollow fiber PVDF membranes. The
experimental setup is shown in Fig. 21.
The RO system was one unit of the direct drinking water preparation system, which
consisted of ozone oxidation, catalyze oxidation and active carbon filtration before
the RO system to get rid of organic compounds. No antiscalants were used. So the
total organic carbon (TOC) analysis revealed that the PRO concentrate contained
less than 1mgL-1
of the TOC. The RO was supplied with the tap water with the
recovery of 50%.
The average concentrations of the major ions were as follows:

48
Ca+2
376.00mg L -1
,
Mg+2
203.00mg L -1
,
Cl-
56.75 mg L -1
,
SiO3
-
11.22 mg L -1
,
HCO3
-
9.25 mmol L -1
,
CO3
-2
0.35 mmol L -1
.
Additional Na2SiO3.9H2O, FeCl2.6H2O and MnCl2.4H2O were added in the PRO
concentrate for special experimental needs.
a. DCMD process of the PRO concentrate
The experimental results shown in Fig. 22 (stage I) demonstrated the direct
application of the PRO concentrate as a feed. At the initial stage of the experiments,
an increase of permeate flux was observed, just as reported in other works. There
were two explanations based on the PP membrane experiments for the initial flux
increase. A significant increase of membrane pores were observed after the water
contact experiments, which lead to the reduction of resistance of the vapor transfer
across the membrane. Another reason was the asymmetrical structure of the
membrane. The pores on the membrane surface were lager, what may lead pores
to be filled with water. It caused a decrease of thickness of the gas diffusion paths
and consequently the DCMD efficiency was increased.
Fig.22, Variation of the efficiency, during the DCMD process, [6]

49
Fig.23, SEM images of the PVDF membranes; (a1) inner surface of fresh membrane (500x); (a2) inner surface of
fresh membrane; (a3) inner surface of fresh membrane after DCMD experiment; (b1) cross section of the fresh
membrane (500x), (b2) cross section of the fresh membrane (2000x); (b3) cross section of the membrane after
DCMD experiment. [6]
The morphology of the PVDF membrane was uniform with the finger-like pores
outside and the sponge structure in the center. Obviously, the pores on the
membrane surface were much smaller than the sponge pores, so the reason of flux
increase was not related to the larger pores on the membrane surface. However, a
slight increase of pore size, difficult to assess visually, can lead the reduction of the
resistances of vapor diffusion across the membrane and the permeate flux may
increase. The permeate conductivity kept decreasing during the process and it
indicated that no wettability phenomenon occurred. The observed initial permeate
flux increase may also result from the dissolved gas transferring across the
membrane when heated.
After the initial increase, a continual decline of the permeate flux was found. A large
amount of deposit was found in the inlet of the membrane module (Fig. 24).

50
Fig.24, SEM micrograph of CaCO3 deposit collected from the inlet of the module, [6]
A concentration decrease of HCO3
-
was found in the feed, which indicated the flux
decline was associated with the formation of CaCO3. So it can be assumed that the
feed flow decrease caused by the deposit clogging was the main reason for the
decrease of the module efficiency. It is because that the feed was pumped with
magnetic pump; therefore, an increase of the feed flow resistance caused a decline
of the flow rate. A decrease of the flow rate caused an unfavorable increase of the
temperature polarization and comprises a possible reason of the observed reduction
of the module efficiency. The feed flow rate can be entirely restored by rinsing the
module with 2% HCl.
The results presented in Fig. 22 (stage II) were obtained after the acidification of the
PRO concentrate. There was a sharp increase of the permeate conductivity at the
beginning of the process. It was resulted from the CO2, which was not degassed
completely after acidification, crossing the membrane to the permeate side. The
problem of CaCO3 was alleviated by acidification, so the permeate flux declined
only 20% after 200 h running after acidification (stage II). Then a sharp flux decline
was observed.
At high levels of water recovery, CaSO4 crystallization may take place when Ca+2
and
SO4
-2
on the feed side exceed the solubility limit of CaSO4. Figure 25 illustrates the
changes of Ca+2
and SO4
-2
during the process, which demonstrates the formation of
CaSO4.

51
Fig. 25, Changes of Ca
+2
and SO4
-2
concentration during DCMD process, [6]
CaCO3 and CaSO4 were prone to precipitate in bulk solutions rather than at
membrane surface, and CaSO4 precipitate can lead to a faster flux decline than
CaCO3 precipitate. This phenomenon can be explained by the different morphology
between them. CaCO3 has a hexagonal structure, and was more tenacious and
compact; while CaSO4 had a needle shape and was loosely attached. So in the
experiment, CaCO3 was found to attach to the feed container and the tubes, the
probability of clogging the membrane module was decreased. While CaSO4 was
found moving freely in the solution, once it formed, it would lead to a rapid clogging
of the module, which would cause a sharp flux decline.
The feed PRO retentate pH was adjusted to 4.0 in stage III. Until the end of the
process, the problem of scaling was eliminated just as reported in other works.
b. DCMD process of silica bearing PRO retentate
Additional Na2SiO3.9H2O was added in PRO retentate, and the initial silica
concentration of PRO retentate was 50 mg L-1
. Permeate flux and conductivity is
shown as a function of elapsed time in Fig. 26.

52
Fig.26, Variation of the efficiency during the DCMD process of silica bearing PRO, [6]
A lot of deposit was found in the inlet of the module. The EDS analysis of the deposit
demonstrated that the deposit contained a large amount of Si, C and Ca with a trace
amount of Mg. The deposit could be assumed as a mixture of CaCO3 and silicate. At
alkaline pH, silica may co-precipitate with soluble metals to form magnesium silicate
(Mg2SiO4) or calcium silicate (Ca2SiO4). Furthermore, silica may be adsorbed onto the
surface of insoluble metal hydroxide compounds, such as Mg(OH)2 or MgCO3.
The observed initial flux decline was due to the impurity clogging the module. After
the washing of the module, the module efficiency recovered to the initial level. The
performance kept stable until the PRO retentate was concentrated about 40 times.
Then the sharp flux decline occurred. A large amount of deposit was found at the
module inlet, while only a little deposit was found on the membrane surface. The
EDS analysis (Fig. 27) showed the deposit was mainly consisted of Si, S, C, Ca, Cl, O
and a trace amount of Cu and Mg. The deposit was assumed as a mixture of calcium
scaling and colloid silica.
Fig.27, SEM-EDS of the deposit collected from the module inlet, [6]

53
There are two relevant categories of silica fouling, namely precipitation fouling and
particulate fouling. Precipitation is a glasslike scale formed on the surface due to
concentration polarization. On the other hand, particulate fouling is the
accumulation of colloids, formed initially in bulk solution or the boundary layer
and then deposit on the membrane surface. The silica fouling potential is
dependent on the concentration of dissolved silica exceeding the amorphous silica
equilibrium solubility of the solution. In RO processes, concentration polarization
phenomenon is significant, so the silica concentration adjacent to the membrane
surface is highly supersaturated; precipitation and particulate fouling were both
common and then the initial silica feed concentration may affect the silica fouling.
However, during the DCMD process, temperature polarization phenomenon plays a
more important role than concentration polarization in the heat and mass transfer
process. So silica fouling phenomenon may be different from the pressure-driven
membrane processes. Precipitation of monomeric silica on the membrane surface
was not found during the process. The SEM analysis showed a mixture of colloid
silica and calcium scaling formed during the process. It seem that particulate fouling
occurred when the bulk silica concentration is high enough to reach the silica
polymerize concentration.
c. DCMD process of iron and manganese bearing PRO retentate
Additional FeCl2.6H2O and MnCl2.4H2O were added to PRO retentate and the initial
iron and manganese concentration of PRO retentate were kept at 1mg L -1
and 0.5
mg L-1
, respectively. The solution pH was adjusted to 6.0. As seen in Fig. 28, the
permeate flux began to decline when the installation run 75 h.
Fig.28, Variation of the efficiency during DCMD process of high iron and manganese PRO retentate. [6]

54
It was found that the major reason of the observed permeate flux decline was the
formation of the deposit at the inlet of the module. The SEM-EDS analysis
demonstrated that the deposit collected from the module inlet was mainly Ca, C and
O, with a smaller amount of Cl. It can be concluded that CaCO3 was still the main
cause of the reduction of membrane efficiency and the influence of iron and
manganese is not significant.
12. RO concentrate treatment with VMD [9]
Simulations were performed to optimise the operating conditions and were
completed by bench-scale experiments using actual RO brines and synthetic
solutions up to a salt concentration of 300 gL-1
. For the membrane studied,
temperature and concentration polarisation were shown to have little effect on
permeate flux. After 6 to 8 h, no organic fouling or biofouling was observed for RO
brines. At high salt concentrations, scaling occurred (mainly due to calcium
precipitation) but had only a limited impact on the permeate flux (24% decrease for
a permeate specific volume of 43 Lm-2
for the highest concentration of salt). Calcium
carbonate and calcium sulphate precipitated first due to their low solubility and
formed mixed crystal deposits on the membrane surface. These phenomena only
occurred on the membrane surface and did not totally cover the pores. The crystals
were easily removed simply by washing the membrane with water. A global
recovery factor of 89% can be obtained by coupling RO and VMD.
Fig.29, Seawater desalination a) by RO conventional process, b) RO+VMD intergrated process, [9]

55
Two points must be carefully explored and will be developed in this study. The first
point concerns the definition of the best operating conditions in order to obtain high
permeate fluxes and minimize the energy consumptions for high salinity solutions.
The second point focuses on the possible limitations of the process by polarisation
effects and/or by fouling.
12.1 Experimental
a. Set-up
All the experiments were performed with a bench-scale batch pilot plant (Fig. 30).
Fig.30, Bench-scale pilot plant, [9]
b. Material and methods
The membrane used in this study was a PTFE flat-sheet. The membrane
characteristics are given in Table 11.
Table 11, Characteristics of PTFE membrane, [9]

56
Fig.31, Computed variation of permeate
flux during the concentration of RO
brines (Pp = 500 Pa, Tf = 50
o
C, Re = 4500)
This membrane has high hydrophobicity, a high Liquid Entry Pressure LEP
(corresponding to the minimal pressure for which liquid is observed to pass) and a
mid-range permeability (determined with the permeate vapour flux measured at
different transmembrane pressures for purified water).
Two types of feed waters were used in this study:
(i) A synthetic concentrated solution containing only the mineral part of seawater at
various total salt concentrations (94.2 g L-1
, 148.6 g L-1
and 291.1 g L-1
respectively for
the SW95, SW150 and SW300 solutions). Its composition is given in Table 12. The
commercial salts were dissolved in purified water.
Table 12, Composition of the synthetic solutions, [9]
(ii) An actual RO brine from an RO plant installed in the Mediterranean Sea. Its
total salt concentration was about 50 g L-1
and its total organic carbon concentration
was between 1.3 and 1.6 mg L-1
. Its conductivity was between 47.8 and 51.4 mS cm-1
at 20o
C.
12.2 Results and discussion
a. Effect of the feed concentration
The objective of the work was to concentrate RO brines from 50 to 300 g L_1.
Simulation was so performed to study the variation of the permeate flux during the
concentration of these brines (Pp = 500 Pa, Tf = 50o
C and Re = 4000). The results are
presented in Fig. 31. As expected, the permeate flux decreases when the bulk salt
concentration increases.

57
The increase in concentration of the solution induces a modification of the
transmembrane pressure difference by decreasing the water vapour partial
pressure of the feed solution. This decrease is due to the decrease in the activity
coefficient when the concentration of the solution increases. It should be pointed
out that, during the concentration of actual brines, a supersaturation of some salts e
which leads to their precipitation and crystallisation be or a fouling can be observed.
b. Effect of the operating conditions: permeate pressure, feed temperature and
Reynolds number
This study aimed to determine, on the basis of the simulation, the best operating
conditions (permeate pressure, feed temperature and hydrodynamics) for highly
concentrated solutions. For each set of experiments, two of the operating
parameters (Pp = 500 Pa, Tf = 50o
C and Re = 4500) were fixed and the other was
varied for the two different membranes and for the three salt concentrations.
Permeate vapour flux was the main criterion but considerations of specific energy
requirements were also helpful for the selection of the operating conditions.
Fig. 32 presents the results of the simulation obtained with the two membranes for a
variation of permeate pressure (Fig. 32a and b), feed temperature (Fig. 32c and d)
and Reynolds number (Fig. 32e and f).

58
Fig.32, Computed variation of permeate flux versus: (a&b) permeate pressure (Tf = 50
o
C, Re = 4500), (c&d) feed
temperature (Pp =500Pa, Re = 4500), and (e&f) Re number (Pp = 500Pa, Tf = 50
o
C). [9]
The trends for the influence of permeate pressure and feed temperature are the
same as previously described for the VMD desalination process with lower
concentration or sodium chloride solution. A low permeate pressure and a high feed
temperature result in a high permeate flux.
When permeate pressure varies from 6100 to 600 Pa (Fig. 32a and b), the
permeate flux is more than doubled whereas the specific energy requirement is
nearly constant, as the energy required to maintain vacuum pressure is only a small
part of the total energy requirement (less than 2%).
Feed temperature is a very sensitive operating parameter, which significantly
influences both permeate flux (Fig. 32c and d) and total energy requirement. It has
a major influence on the water vapour partial pressure (according to the exponential

59
Antoine Equation) and thus on the transmembrane pressure difference. For
example, increasing temperature from 20 to 70o
C increases the permeate vapour
flux from 1.33 to 22.30 L h-1
m-2
(for SW300 and KM = 3.26 10-6
smol1/2
m-1
kg-1/2
) but
the total energy requirement is drastically increased too. At 70o
C, more than 99% of
the total energy requirement is heat energy.
The effect of hydrodynamics on the permeate flux is strongly dependent on the
Knudsen permeability of the membrane (Fig. 32e and f). From a laminar flow (Re =
350) to a turbulent flow (Re = 6100), the permeate flux increases by 13-15% for the
low-permeability membrane (Fig. 32e) and by 44-50% for the more permeable
membrane (Fig. 32f), whatever the solution concentration. For the low-permeability
membrane, the temperature and concentration polarisation have been shown to
have little effect, which is confirmed by the independence between the permeate
flux and hydrodynamics. However, when the membrane permeability is higher, the
permeate flux is higher and so concentration and temperature polarisation might
be higher, depending on the Knudsen permeability value and Re value. Increasing
turbulence of the flow will increase heat and mass transfer coefficients in the
boundary layer near the membrane. The permeate flux approaches an asymptotic
value when Re increases. Beyond this Re value, hydrodynamics has less effect on the
permeate flux. Although the specific energy requirement for feed circulation is only a
small part of the total energy requirement, it does not seem necessary to operate at
too high Re. A flow with a low level of turbulence (Re = 4500) could be a good
compromise.
c. Effect of the membrane permeability
As shown previously, the membrane permeability is a key parameter for the VMD
performance. Indeed, with no apparent increase in the energy requirement, the
permeate flux increases dramatically when the membrane permeability is increased.
Fig. 33 shows the effect of the Knudsen permeability on the permeate flux for a set
of operating conditions selected for their ability to allow a high permeate flux and to
minimize energy requirements (Pp = 500 Pa, Tf = 50o
C and Re = 4500).

60
Fig.33, Computed variation of permeate flux versus Knudsen permeability for different solutions
concentrations, (Pp = 500 Pa, Tf = 50
o
C and Re = 4500). [9]
As expected according to results obtained with “normal” seawaters, for highly salty
solutions, higher fluxes are attainable with high Knudsen permeability membranes.
Knudsen permeability depends on membrane structure. High Knudsen permeability
can be obtained by using a thinner membrane, or a membrane with a larger
number of pores and/or larger pores. However, even though a slight increase in
Knudsen permeability value dramatically improves water permeate vapour flux, it
may also decrease membrane hydrophobicity and no longer prevent water in the
liquid phase from passing through membrane pores. Wetting is a sensitive point in
membrane distillation. A balance must be found between the Knudsen permeability
value and membrane hydrophobicity.
12.3 VMD performance with concentrated synthetic brines
As mentioned previously, the model used for the simulations does not take any
progressive fouling (scaling, organic fouling or biofouling) into consideration. In order
to investigate whether these phenomena were limiting, experiments with different
kinds of concentrated solutions were performed. The next part of the paper will
focus on experiments performed with the Fluoropore membrane (KM = 3.26 10-6
smol1/2
m-1
kg-1/2
).

61
Table 3 shows the operating conditions for the different sets of experiments (A-D):
Tf: inlet feed temperature
Pp: permeate pressure
Re: feed Reynolds number
Ji: initial experimental permeate flux for the solution studied
Jth: initial permeate flux calculated by modelling
Table 13, Initial experiment operation conditions and permeate fluxes, [9]
As a preliminary study, the influence of temperature and concentration polarisation
on the permeate flux and on the salt saturation for the different sets of experiments
was estimated.
a. Effect of the radial polarisations on the permeate flux
The feed radial profile or so called polarisation is linked to the boundary layer close
to the membrane surface. Close to the membrane, the temperature Tm is lower than
the feed bulk temperature Tf whereas feed salt concentration Cm is higher than the
feed salt concentration Cf in the bulk. These temperature and concentration
polarisations can influence the mass and heat transfer by introducing a new thermal
and/ or mass resistance. Different polarisation ratios were calculated using the VMD
model.
Table 14 presents radial temperature and concentration polarisations and their
effects on the flux. Tm/Tf, named the “temperature polarisation coefficient”, is the
ratio between the temperature close to the membrane calculated by simulations
and the experimental temperature measured at the feed side. Cf/Cm, named the
“concentration polarisation coefficient”, is the ratio between the concentration in
the feedwater and the concentration close to the membrane calculated by
simulations.

62
Table 14, Radial concentration and temperature polarisations, [9]
The temperature polarisation coefficient (Tm/Tf) was nearly equal to 1 whereas the
concentration polarisation coefficient (Cf/Cm) lay between 0.970 and 0.987. The
concentration was slightly higher at the membrane surface than in the feed bulk.
This polarisation of concentrations seemed to be a little greater for high salt
concentration. Concentration polarisation was a littlemore marked than
temperature polarisation.
Jm(Tm, Cm)/Jf is the ratio between the permeate flux calculated at Tm and Cm i.e. in the
conditions close to the membrane, and Jf. This ratio represents the permeate flux
polarisation coefficient taking both temperature and concentration polarisation into
consideration. It is the global permeate flux reduction due to both polarisations.
Jm(Tm, Cf)/Jf is the ratio between the permeate flux calculated at Tm and Cf, and Jf.
This ratio represents the permeate flux polarisation coefficient considering only the
temperature polarisation. It is the permeate flux reduction due to temperature
polarisation.
Jm(Tf, Cm)/Jf is the ratio between the permeate flux calculated at Tf and Cm, and Jf.
This ratio represents the permeate flux polarisation coefficient considering only the
concentration polarisation. It is the permeate flux reduction due to concentration
polarisation.
Table 15 recalls the global permeate flux reduction Jm(Tm, Cm)/Jf. It also shows the
contribution of temperature polarisation and of concentration polarisation to this
permeate flux reduction. The contribution of temperature polarisation was
calculated by comparison of the permeate flux polarisation coefficient due to
temperature polarisation Jm(Tm, Cf)/Jf and the global permeate flux reduction Jm(Tm,
Cm)/Jf . The contribution of concentration was calculated in the same way.
Table 15, Radial concentration and temperature polarisation, [9]

63
Jm(Tf,Cm)/Jf (Table 4) is always nearly equal to 1, which means that the
concentration polarisation has little effect on the permeate flux. Jm(Tm, Cf)/Jf is
always lower and close to 0.95-0.97. Temperature polarisation has a greater effect
on the flux reduction. This is confirmed by calculation of both contributions (Table
15): the contribution of temperature polarisation to the flux reduction represents
more than 88%for SW95, SW150 and RO brines. Indeed, temperature modification
affects the vapour partial pressure which is an exponential function according to
Antoine’s Law, whereas the concentration modification affects the activity
coefficient, which is less sensitive to variations. For the highly concentrated solution
(SW300) the contribution of temperature polarisation is reduced (62%) and the
effect of concentration polarisation on the flux reduction then becomes more
significant.
In all cases, global permeate flux reduction due to polarisations Jm(Tm, Cm)/Jf is
always higher than 0.94: temperature and concentration polarisation exists in VMD
for highly concentrated seawaters but can be considered to have little effect in
terms of flux reduction, for the membrane and the operating conditions of our
study.
b. Effect of the radial polarisations on salt saturation
Although temperature and concentration polarisation seem to have a limited
impact on the permeate flux, they may play a role in the precipitation of salts.
Calcium carbonate and calcium sulphate are the predominant sparingly soluble
salts present in seawater. Previous works have shown that CaCO3 mainly
precipitates as calcite and CaSO4 as gypsum: CaSO4, 2H2O. These two polymorph
forms were therefore considered in the present work, together with halite as NaCl is
the major compound in seawater. As a measure of the scaling potential of these
salts, the saturation index (SI) was calculated.
The saturation index SIf was calculated for the different experiments in the initial
feed conditions. Results are presented in Table 16.
Table 16, Saturation indices of the salts, [9]
If the saturation index is equal to 1, the solution is at saturation. If the saturation
index is lower than 1, salt will not precipitate. If the saturation index is higher than 1,
salt will be at supersaturation and can precipitate. It should be noted that, in some
cases, precipitation can occur before saturation.

64
Fig.34, a)Normalized permeate flux and b) Apparent Knudsen permeability versus specific volume for
SW 95 (a, Table 13), SW150 (b, Table 13) and SW300 (c, Table 13). [9]
According to these results, the precipitation of NaCl should not be a problem in any
case. Precipitation of gypsum can occur for the solutions with concentrations higher
than 150 g L-1
. The main problem is linked with the precipitation of calcite.
Solubility of CaCO3 and CaSO4 in water decreases with temperature. The
temperature is slightly lower at the membrane due to temperature polarisation but
the salt concentration is higher. Temperature polarisation and concentration
polarisation thus have opposite effects on the possible precipitation of these salts.
Table 17 shows the ratio between the saturation index calculated at the bulk
conditions (SIf) and calculated at the membrane conditions (SIm) for both calcite and
gypsum. All the ratios are lower than 1, which means that the saturation index is
always higher at the membrane and that precipitation is more likely to occur close to
the membrane.
Table 17, Saturation indices at the membrane and in the bulk, [9]
At normal pH, scaling will obviously occur during the VMD desalination of RO
brines. This scaling is enhanced by the polarisation effects since precipitation is
more likely to occur close to the membrane.
c. Study of the time variation of permeate flux for synthetic brines: evaluation of
scaling effects
In order to study the possible effects of scaling on flux, variations of permeate flux
for experiments of 6-8 h are presented in Fig. 34 for the SW95 (Experiment A from
Table 13), SW150 (Experiment B from Table 3) and SW300 solutions (Experiment C
from Table 13).

65
Fig. 6a presents the values of normalised flux versus the permeate specific volume
i.e. the permeate volume transferred per unit membrane area. Values of flux are
normalised to the initial permeate flux i.e. the permeate flux obtained for each
solution at the beginning of the experiment (Ji in Table 13).
Table 18 reports the initial permeate flux and the permeate flux obtained for a
permeate specific volume of 43 L m-2
. The permeate flux decline was 11%, 8% and
24% for SW95, SW150 and SW300 respectively.
Table 18, Values of permeate flux for different experiments. [9]
All the experiments were performed in a batch reactor. The decrease in the
permeate flux can be explained in two ways. Firstly, it may be due to the
modification of the feed water vapour partial pressure caused by the
concentration of the feed solution linked to the permeate volume filtered as
increasing seawater concentration decreases the water molar fraction Xwater and the
feed water activity coefficient awater. The permeate flux Jwater is thus reduced when
the transmembrane pressure difference decreases. Secondly, it may be due to
fouling on the membrane surface.
In order to isolate these two phenomena, a new Knudsen permeability, called the
apparent Knudsen permeability, was calculated versus time:
Calculation of this apparent Knudsen permeability was based on some experimental
data with some assumptions validated previously in this paper:
- Tm is equal to the feed bulk temperature Tf (no temperature polarisation effects)
- Xwater close to the membrane surface is equal to Xwater in the bulk (no concentration
polarisation effects).
The apparent Knudsen permeability represents the permeability of the membrane
during the experiment. This permeability takes possible modification of membrane
properties by fouling (reduction of the surface porosity, reduction of pore size.) into
consideration. Fig. 34b shows the apparent Knudsen permeability versus the
permeate specific volume. After a filtration of 43 L m-2
, the decline of the apparent
Knudsen permeability was calculated. The decline was 8%, 9% and 18% for the
SW95, SW150 and SW300 respectively. The Knudsen permeability variations were

66
very close to those of the permeate flux, which seems to indicate that the decrease
of the permeate flux is only linked to the modification of the apparent Knudsen
permeability. The effect of the concentration can be neglected for these
experiments. Nevertheless, it should be noted that the decline of the permeate flux
was greater than the decline of the Knudsen permeability for the SW300
experiment. As already shown (Table 14), effects of concentration are greater when
the concentration is higher. Finally, the flux decline seems very limited in all cases.
12.4 Observation and study of scaling
The observed modification of the Knudsen permeability may be linked to a
modification of membrane properties by scaling. Observations of the fouled
membrane (after drying) with SEM and analysis with EDS probe allowed the
membrane surface to be visualised and some deposited components to be
identified. Precipitation may also have occurred during the drying of membrane but
this point was not considered here. Fig. 35 presents a view of an unused membrane
and membranes after permeation of SW150 and SW300 solutions.
Fig.35, SEM micrograph of a) an unused membrane (x1500), b) a membrane after the SW 150 experiment
(x1500) and c) a membrane after the SW300 experiment (x1500). [9]

67
Fig.35, EDS analysis of
the membrane after the
SW150 experiment,
showing a) free
membrane surface, b) a
calcium crystal and c) a
mixed zone. [9]
a
b
Fig. 35b shows the membrane surface after the experiments with SW150. Three
areas can be distinguished: a free membrane surface, an area with isolated crystals
and an area with a mixture of crystals. These different zones were analysed by EDS
probe (Fig. 36).

68
The first area (Fig. 35a) is a free membrane zone. On this area, pores and the
membrane surface are visibly free. Only traces of calcium can be found. This large
clear area of the membrane explains why the permeate flux decline caused by
scaling (8%) was not too great. The second area (Fig. 35b) shows some isolated
crystals. These are calcium crystals with either sulphate or carbonate. Finally, the
third area (Fig. 35c) is an area composed of many different crystals in large
quantities, with the presence of: calcium, oxygen, sodium, potassium, chloride,
sulphur, bromide, magnesium, and carbon. It seems that all the salts present in the
feed water have precipitated in this area. Using the presence of these three zones,
hypotheses can be formulated for the crystallisation mechanisms. Calcium crystals
seem to be the first to precipitate, initially as a trace on the membrane surface,
then as crystals associated with other ions. This is easily explained by the low
solubility of calcium salts. These first crystals can be nuclei for the precipitation of
the other salts. Then a mixture of different salts is formed and it progressively covers
the membrane surface. Fig.36 summarises the stages for the precipitation on the
membrane surface.
Fig.36, Schematic represantation of precipitation stages on the membrane. [9]
12.5 Study of scaling for SW300 solution, the highest feed concentration solution
Fig. 8c shows the presence of more and more salts when the concentration increases
for SW300 (Experiment C from Fig. 6). However, the pores are always partially free
and the scaling is not particularly marked in comparison with the SW150 experiment.
Since the beginning of the experiments, the SW300 solution had shown some
precipitation in the feedwater. If the precipitation had already occurred in the feed
water, crystals may have been used as nuclei for the other precipitations.
Precipitation seems to have occurred both on the membrane surface and in the
feedwater, thus preventing the precipitation from being too high on the membrane
surface.

69
12.5 Conclusion on scaling in VMD of concentrated synthetic brines
The study of scaling has shown that membrane scaling obviously exists in VMD for
desalination. It is greater at high salt concentration and seems to occur on the
membrane surface. The main element responsible for scaling is calcium. However,
on the time scale of our experiments, the scaling did not have a very marked effect
on the permeate flux obtained. Moreover, for the different concentrations of
synthetic feed water, the permeability was measured after washing the membrane
with osmosis water. No significant difference was observed between the initial
permeability and the permeability after washing. This confirms that the scaling
occurred only on the surface and was reversible. The next part will focus on the
membrane distillation of an actual RO brine in order to determine the possible
contribution of organic fouling and biofouling.
12.6 Membrane distillation of actual RO brines
A five-day experiment was performed with actual RO brines. Fig. 10 presents the
variation of the normalised permeate flux and of the apparent Knudsen permeability
versus the permeate specific volume.
Table 19, Results of simulation of the coupling of RO and VMD. [9]
A decrease of the permeate flux was observed in the first two days, linked to a
decrease of the apparent permeability. However, after these first two days, the
permeate flux and the apparent permeability remained constant. This seems to
indicate that there was no effect of scaling, organic fouling or biofouling. It should be
noted that the salt concentration (about 50 g L-1
) was much lower than in the
synthetic solutions and the concentration of organic matter (1.3 and 1.6 mg L-1
of
TOC) was very low compared to the total salt concentration.
Only very few crystal deposits were observed on the membrane surface: a mixture of
the different crystals previously observed (calcium sulphate). However, these
deposits did not cover the pores.

70
12.7 Conclusion
On the basis of simulation and experiments at bench-scale, VMD has proved to be
very interesting when integrated in a seawater treatment line as a complementary
process to SWRO.
- Thanks to the coupling of SWRO and VMD, high recovery can be obtained (89% in
this study), corresponding to a brine volume reduction by a 5.5 factor, and the water
production can be increased by more than 2.
- With correctly chosen operating conditions (low permeate pressure, high feed
temperature, turbulent fluid regime), even for a membrane with medium
permeability, high permeate fluxes can be obtained. For a permeate pressure of
6000 Pa, a temperature of 50 o
C and a Re of 4000, permeate flux ranges from 17 L h-1
m-2
to 7 L h-1
m-2
during the concentration of RO brines from 64 g L-1
to 300 g L-1
- Temperature and concentration polarisation have little effect on the permeate flux
even for the high salt concentrations.
- For high salt concentrations, scaling occurs in vacuum membrane distillation but its
impact on the permeate flux is very limited. Large areas of membrane remain free of
visible fouling. The main salts responsible for the scaling are calcium crystals such as
calcium carbonate and calcium sulphate, which have the lowest solubility. Mixtures
of different crystals are also often found on the membrane surface.
- Nevertheless, in all the cases, scaling is only reversible, surface scaling that can be
easily removed by a simple washing of the membrane with water.
- The concentration of organic matter is too low to show an impact on permeate flux
at the time scale of a few hours or days but its impact on membrane hydrophobicity
and wetting for long-term operation must be studied carefully.
13. Integration of accelerated precipitation softening with MD for PRO
concentrate [8]
Currently, the available method to improve RO recovery is using water soluble
polymeric antiscalants. It can suppress mineral salt precipitation to some extent.
However, even with the use of antiscalants, mineral salt scaling remains an
impediment to achieving high product water recovery, partially due to the increased
potential of fouling when excessive dose of antiscalants is applied.
Another possible approach to improve RO recovery is to utilize an accelerated
precipitation process to remove scaling ions before RO process. Conventional
precipitation is to induce calcium carbonate crystallization through chemical dosing
(e.g. lime, caustic, and soda ash). However, the produced calcium carbonate crystals

71
require a long time (about 1.5–3 h) to settle and result in low content solid (2–30%).
Due to the above drawbacks, seeded precipitation has been developed. It can
provide a preferential surface area for heterogeneous nucleation and growth of
mineral salts. Thus, the precipitation kinetics and the efficiency of solid–liquid
separation can be significantly improved. Various seeded precipitation softening for
RO process pretreatment were reported in the literature.
In the present work, a hydrophobic membrane process-DCMD was integrated with
APS for high-recovery desalting of PRO concentrate. The integrated process involved
inducing and accelerating mineral precipitation by sodium hydroxide dosing,
followed by solid–liquid separation, microfiltration and subsequent the DCMD
desalting.
The ΑPS process, between the PRO concentrate and the DCMD process, involved pH
adjustment with sodium hydroxide along with calcite seeding, followed by
microfiltration to avoid seeds clogging of the DCMD module. Elemental analysis
revealed that APS treatment enabled 92% removal of calcium, 58.4% removal of
total hardness, 4.4% removal of magnesium, 1.1% removal of sulfate and 1.6%
removal of silica. Compared with the sharp decline found in the DCMD process of the
PRO concentrate, the permeate flux declined only 20% within 300 h running after
APS treatment. Then the PRO concentrate was concentrated 40 times and the whole
recovery was enhanced to 98.8%.
Accordingly, the objective of this work were to (1) evaluate the sodium dosing, seed
dosage and agitating rate via small scale calcium removal tests; (2) demonstrate the
integration of APS with DCMD at a laboratory-scale to achieve high water recovery;
and (3) evaluate the hydrophobic stability of the PVDF membrane via the variation of
the permeate flux and conductivity during 300 h continual running.
13.1. PRO concentrate and reagents
The quality of PRO concentrate was shown in Table 20. Calcium carbonate powder
(>99.0%, A.C.S, Reagent, 10 μm) and quartz sand (powder, 250 mesh) were used as
calcite and quarts sand seeds, respectively. All the other reagents were analytical
grade.

72
Table 20, Quality of the PRO concentrate, [8]
13.2 APS-DCMD process
The laboratory-scale APS-DCMD setup is illustrated in Fig. 37. The APS treatment
with actual PRO concentrate was conducted in a 10 L crystallizer. This apparatus
consisted of a crystallization reactor with a conical bottom, and a poleless speed-
adjusting agitator.
Fig.37, ΑPS-DCMD process set-up,(1) 10L crystallization reactor,(2) magnetic pump, (3) cartridge filter, (4)
magnetic pump, (5) feed reservoir, (6) membrane module, (7) cooling oil, (8) permeate reservoir, (9)
thermometer, (10) flow meter and (11) conductivity monitor. [8]
13.3 Membrane module
The DCMD membrane module was made by a polyester tube and two UPVC T-tubes.
The module was equipped with 50 hydrophobic self-made hollow fiber PVDF
membranes.

73
13.4. Determination of the optimum softening conditions
13.4.1 Initial pH
Fig. 38 shows the calcium concentration decline at different initial pH without seeds.
The decline expressed as the ratio (Ca+2
)/(Ca+2
)0,
Fig.38, Precipitation kinetics for PRO concentrate at different PH without seeds. [8]
where (Ca+2
) is the calcium concentration at time t and (Ca+2
)0 is the initial calcium
concentration. It can be noted that, in the absence of NaOH dosing (pH 7.70),
calcium concentration of the PRO concentrate decreased by about 10% relative to
the initial calcium concentration, over a period of 1.5 h. However, once the PRO
concentrate was dozed with NaOH, the calcium concentration declined rapidly. As
the solution pH increased from 9.10 to 11.10, the calcium removal efficiency
increased from 38% to 86%, over a period of 30 min. It is due to the deprotonation
of bicarbonate ions that the concentration of CO3
-2
was generated at higher pH. This
resulted in a higher initial calcium carbonated supersaturation and thus higher
degree of CaCO3 precipitation. Also, higher initial calcium carbonate supersaturation
resulted in faster precipitation kinetics as indicated by the greater rate of calcium ion
depletion.
It should be recognized that Mg(OH)2 would form when solution pH achieved >10.50
according to the saturation index. However, Mg(OH)2 precipitation is a gel-like
structure with high capacity for water retention t. Therefore, in water softening
processes, one typically avoids operating in the range of excessive Mg(OH)2
precipitation due to difficulties in solid-liquid separation and solid dewatering
operations. Therefore, initial pH adjustment to about 10.10, for APS treatment, was
selected for evaluating for the feasibility of attaining high product water recovery.

74
13.4.2. Seed selection and dosage
The impact of calcite seed and quartz sand on precipitation softening was initially
carried out with 800mL PRO concentrate. The former experiments (Fig. 39) showed
in the absence of seeding, calcium removal about 86% was attained for pH 10.10
within 30 min.
Fig.39, Precipitation kinetics for PRO concentrate using different seeds and different calcite dosage. [8]
It can be noted from Fig. 39 that 10 g/L quartz sand and calcite resulted in even
faster precipitation kinetics, within a period time of 15min. After steady-state was
attained, 10 g/L quartz sand result in 87% calcium removal while calcite lead to an
even higher removal efficiency, as high as 94%.
Calcium removal ration as a function of elapsed time at lower calcite seed load (3
and 5 g/L) is also showed in Fig. 39. The steady-state was approached within 15 min
for the three calcite seed load. The results showed 3 g/L calcite seed load resulted in
87% removal of calcite, and 5 and 10 g/L seed load resulted in 91% calcite removal
and 94% calcite removal, respectively. It is noted that operating at lower calcite
seed load (5 g/L) can also yield high calcium removal. It suggested that sufficient
surface area for precipitation was achieved at this lower seed loading, so 5 g/L was
chosen to be the optimal calcite seed load.

75
13.5. Performance of integrated APS-DCMD process
Following the results of the small-scale calcium tests, a large scale APS treatment of
PRO concentrate was conducted in a 10 L crystallizer at pH 10.10 and calcite seed
load of 5 g/L. Elemental analysis (Table 21) of the PRO concentrate before APS and
after APS revealed that, calcium decreased from 118.20 to 9.72mg/L, with the
removal efficiency about 92%.
Table 21, Quality of the PRO concentrate before and after APS treatment. [8]
In addition, the analysis indicated 4.4% removal of magnesium, 1.1% removal of
sulfate and 1.6% removal of silica. There was also a measurable 58.4% removal of
total hardness, which was mainly caused by the calcium depletion.
Fig. 40 illustrates the DCMD performance of the PRO concentrate before and after
APS treatment.
Fig.40, Variation of permeate flux and conductivity as a function of elapsed time. [8]
In Fig. 40 stage I, a slight increase of permeate flux was observed at the initial period
of the process, just as reported in other works. After the increase, a continual
decline of the permeate flux was found. SEM analysis of the module showed that a
certain amount of deposit was found at the inlet of the module (Fig. 41(a)),

76
Fig.41, SEM images, (a) Deposit found at the inlet after 50h DCMD performance before APS; (b) deposit found
at the inlet after 300h DCMD performance after APS and (c) inner surface after 300h DCMD performance after
APS. [8]
however, little deposit was found at the membrane inner surface. So it can be
inferred that the deposit found at the inlet of the membrane module was the
major reason of the permeate flux decline. That is because the feed was pumped
with magnetic pump, the deposit may cause a clogging of the feed flow channel
which leads to a feed flow resistance, and therefore, cause a decline of the flow rate.
A feed flow rate decrease from 0.60 to 0.45 m/s was observed, and this decrease
caused an unfavorable increase of the temperature polarization and comprises a
possible reason of the observed reduction of the module efficiency. The EDS analysis
showed the deposit was mainly consisted of Ca, C and O, and also a concentration
decrease of HCO3
−
was found in the feed, thus it can be concluded the deposit was
mainly CaCO3. The feed flow rate can entirely restored by rinsing the module with
2% HCl. The permeate conductivity kept decreasing and stabilized at about 3.5
μS/cm at last, which indicated that the PVDF membrane exhibited a stable
hydrophobicity.
Fig.41, stage II presents DCMD performance of the PRO concentrate after APS
treatment. The elemental analysis (Table 21) showed that, after the APS treatment,
the calcium concentration had such a significant decrease that the probability of
CaCO3 and CaSO4 scaling was significantly decreased. Thus permeate flux declined
only 20% after 300 h running, then the PRO concentrate was concentrated 40 times
and the whole recovery was enhanced to 98.8%. A little deposit was found at the
inlet of the module (Fig. 41(b)), and the membrane inner surface was also covered
with a little deposit (Fig. 41(c)). The EDS analysis showed that the deposit found at
the inlet of the module and the membrane surface was both consisted of Mg, Si, C, O
and Si.
The permeate conductivity kept from 2.0 to 4.0_μS/cm during of the process, and
slightly increased to 6.0μS/cm at the end. That may be associated with partial
wetting phenomenon caused by large pores just as mentioned in other works. Large
pores inevitably exist in the membrane and lead to a low LEPw. LEPw is the minimum
pressure at which water will overcome the hydrophobic forces of the membrane and

77
will permeate pores. In an integrated NF/MD demineralization process, after 90h
DCMD performance, the permeate conductivity was found increasing from 1.5 to
2.0μS/cm. The author inferred the phenomenon was associated with partial
wettability caused by 1μm large pores. Gryta pointed that a deterioration of
permeate quality might occur when a partial wetting phenomenon will take place,
and in the case when a wetted membrane area is not to large, MD process may be
still continued. However, the influence of large pores may become serious at higher
concentration just like the slight permeate conductivity increase at the end of the
DCMD process.
13.6 Conclusions
In the present work, accelerated precipitation softening was integrated with direct
contact membrane distillation for high recovery desalination of primary reverse
osmosis concentrate. The optical solution pH, calcite dosage and agitation rate for
APS were 10.10, 5 g/L and 200 r/min, respectively. Experimental results indicated
that APS treatment enabled up to 92% removal of calcium, thus, CaCO3 and CaSO4
scaling was alleviated during the DCMD process. It can be noted that permeate flux
declined only 20% within 300 h running, then the PRO concentrate was concentrated
40 times, and the whole recovery was enhanced to 98.8%.
14. Sustainable operation of MD for mineral recovery from hypersaline
solutions [7]
Direct contact MD (DCMD) experiments were performed with water from the Great
Salt Lake (4150,000 mg/L total dissolved solids) as the feed stream and deionized
water as the distillate stream. DCMD was able to concentrate the feed solution to
twice its original concentration, achieving close to complete inorganic salt rejection.
During experiments water flux declined to 80% of its initial value (real-time
microscopy revealed that precipitation of salts on the membrane surface was the
main contributor to the decline in water flux. The application of novel scale-
mitigation techniques was highly effective in preventing scale formation on
membrane surfaces, sustaining high water flux and salt rejection, and eliminating
chemical consumption used for membrane cleaning. MD was compared to natural
evaporation and was found to potentially replace 4047m2
(1 acre) of evaporation
ponds with approximately 24m2
of membrane area and to be nearly 170 times faster
in concentrating hypersaline brines.
In mineral production, evaporation ponds are traditionally utilized for concentration
of saline water and precipitation of minerals, which are then further processed in
chemical plants. Evaporation ponds commonly use large areas, they are time and
energy intensive, and when used, large volumes of valuable water are lost to the
atmosphere. In order to improve the efficiency of mineral recovery, replacement of

78
evaporation ponds with desalination processes could minimize land use and increase
water recovery from hypersaline streams.
14.1. DCMD for concentration of super saturated solutions in mineral production
Replacing traditional concentration methods with MD could produce high quality
minerals and water, reduce land footprint of evaporation ponds, and eliminate the
required pumping of water from pond to pond in mineral production sites. Recent
studies have shown that MD consumes less energy than traditional thermal
distillation such as multi-stage flash and multi-effect distillation, and can further
concentrate brines from desalination processes such as RO, NF and ED. Furthermore,
utilization of low-grade heat sources such as industrial heat emissions and solar
energy can offset the overall energy consumption needed for MD. Recent studies
have coupled membrane processes with crystallizers to concentrate and recover
minerals in hyper saline solutions; however, none of these studies have effectively
mitigated membrane scaling. While membrane scaling has been investigated,
effective scale mitigation techniques for maintaining and restoring water flux and
salt rejection when desalinating saturated solutions are still lacking. In the current
study, DCMD was applied to concentrate Great Salt Lake (GSL) water. The main
objectives of the study were to evaluate the performance of DCMD in concentrating
hypersaline brines from the GSL and in doing so, optimize operating conditions to
maximize water recovery and mitigate membrane scaling. Several unique methods
were developed and tested to identify and mitigate membrane scaling. Finally, the
replacement of evaporation ponds with DCMD was assessed as a means to intensify
the mineral production process.
14.2. Materials and methods
14.2.1. Membranes
Two hydrophobic microporous membranes were acquired from GE Water. The first
membrane (TS22) is a composite membrane consisting of a thin
polytetrafluoroethylene (PTFE) active layer and a polypropylene woven support
layer. The second membrane (PP22) is an isotropic membrane made of
polypropylene (PP).
14.2.2. Bench-scale system
Bench-scale experiments were performed to investigate water flux, salt rejection,
and membrane scaling. A flow schematic of the test unit is illustrated in Fig.1.

79
Fig.42, Flow schematic of the DCMD bench scale system. [7]
Table 22, Ionic composition of the GLS water. All values are for the cartridge filtered GLS water. [7]

80
14.3. Results and discussion
14.3.1. Pure water permeability experiments
Water flux as a function of feed temperature for the two membranes is shown in
Fig.43.
Fig. 43, Water flux as a function of feed temperature for experiments performed with the TS22 and PP22 MD
membrane. The distillate temperature (Td) was either (a) 20 or (b) 30
o
C ,the feed temperature was 30-70
o
C
and the flow rates were kept constant at 1,6 Lxmin
-1
. [7]
Water flux increased exponentially with increasing temperature difference (or
vapor pressure driving force) across the membrane. The water flux through the
TS22 was consistently higher than the water flux through the PP22. This is because
the PP22 is a thicker membrane with more tortuous pores that increase the
resistance to vapor diffusion through the membrane pores, thus resulting in a lower
permeability. Based on these results, a high temperature differential (ΔT=40o
C) and
low temperature differential (ΔT=20o
C) were chosen for the successive batch
concentration experiments; the temperature of the distillate stream was kept at
30o
C and the temperature of the feed stream was either 50 or70o
C.
14.4 Direct contact membrane distillation batch experiments
14.4.1. Successive batch experiments: water flux and salt rejection
Water flux as a function of GSL water total solids concentration for experiments
performed with the two membranes operate data ΔT of 40 and 20o
C is shown in
Fig.44.

81
Fig.44, Water flux as a function of total solids concentration for the successive batch experiments (runs)
performed with (a) the TS22 membrane and (b) the PP22 membrane. Experiments were performed with
filtered GLS water as feed and deionized water as distillate. The distillate temperature was 30
o
C and the feed
temperature was either 50
o
C (50/30) or 70
o
C (70/30). The flow rate was kept constant at 1,6 Lxmin
-1
. [7]
In all experiments the water flux gradually declined as the feed solution
concentration increased, and thus the partial vapor pressure of water in the feed
solution decreased. Thereafter, a sharp decline in water flux was observed in all
the experiments.
Compared to the pure water permeability experiments, a lower initial water flux was
observed during batch experiments performed with the GSL water. For example, in
experiments performed with the PP22 and GSL feed solution, the initial water flux
was 11% (ΔT of 20o
C) and 20% (ΔT of 40o
C) lower than that of the pure water
permeability experiments performed with the same membrane and the same
temperature differences. The experiments performed with the TS22 and GSL feed
solution resulted in an initial water flux of 28% (ΔT of 20o
C) and 38% (ΔT of 40o
C)
lower than in the pure water permeability experiments performed with the same
membrane and temperature differences. The percent decrease in initial water flux
for the experiments performed with the GSL feedwater was lower for temperature
difference of 20o
C and for the PP22. The lower partial vapor pressure of water in
the highly concentrated feed solution (150,000mg/L total solids) was the main
reason for a lower driving force across the membrane and the lower initial water
flux.

82
At increased temperature differentials (ΔT of 40o
C), a higher terminal GSL water
concentration was achieved with the PP22. Although experiments conducted with
the TS22 at a higher temperature difference (ΔT of 40o
C) resulted in greater initial
water fluxes compared to those with the PP22, the initial water flux decreased after
each successive batch experiment. Several studies have reported that polarization
effects are more severe at increased operating temperatures; however, these
effects are reduced when using membranes with low thermal conductivities. Thus,
it is likely that the lower thermal conductivity of the thicker PP22 mitigated the
effects of temperature polarization when tested at increased temperature
differences across the membrane. Temperature polarization effects were more
severe during experiments performed with the TS22 at increased temperature
differences, which resulted in increased membrane scaling. Higher flux also results
in higher heat fluxes, which subsequently decreases the temperature difference
across the membrane and increases temperature polarization effects.
Temperature polarization also affects the mechanisms of membrane scaling.
Because the temperature is lower at the feed–membrane interface than in the bulk
solution, the calcium species solubility is expected to increase, whereas NaCl
solubility is expected to decrease. Additionally, increased water flux results in
increased rates of scaling. Therefore, it is expected that the potential for scaling of
NaCl will be higher for the TS22 membrane (higher flux) than the PP22 membrane.
Water flux and distillate conductivity as a function of time are shown in Fig. 45 from
data presented in Fig. 44.
Fig.44, Water flux and distillate conductivity as a function of total solids concentration for the successive batch
experiments (runs) performed with (a) the TS22 membrane and (b) the PP22 membrane. Successive batch
experiments were performed with filtered GLS water as feed and deionized water as distillate. The distillate

83
temperature was 30
o
C and the feed temperature was either 50
o
C (50/30) or 70
o
C (70/30). The flow rate was
kept constant at 1,6 Lxmin
-1
. [7]
The distillate conductivity for both sets of experiments performed at a ΔT of 20o
C
continuously decreased, indicating that the membrane rejected nearly 100% of all
inorganic and non-volatile constituents. However, the distillate conductivity
increased in both experiments operated at a higher temperature gradient. The
distillate conductivity increased throughout the experiment performed with the
PP22. In experiments performed with the TS22, the distillate conductivity started to
increase only towards the end of the experiment, indicating that the TS22 is less
susceptible to wetting at higher operating temperatures. A previous study by Gryta
reported that operating at temperatures greater than 68o
C may reduce the
hydrophobicity of polypropylene membranes and lead to membrane wetting.
Saffarinietal reported that PTFE membrane support layers showed no signs of
degradation when exposed to high temperatures (<350o
C); however, the potential
for wetting of PTFE membranes does increase with increasing feed temperatures
and salinities.
When operated at a higher ΔT of 40o
C, the water flux had more than tripled,
increasing the overall water flux from 12.8 to 47 Lm-2
h-1
during experiments with the
TS22 and from 12.3 to 40 Lm-2
h-1
during experiments with the PP22. Yet, while the
process was accelerated at high temperature differences across the membrane, both
salt rejection and initial water flux decreased after each successive batch
experiment. When experiments were conducted at a ΔT of 20o
C, the salt rejection of
the membrane was high, the water flux after each successive batch run was
restored, and the GSL water was concentrated to more than 300,000mg/L total
solids. Therefore, operating temperatures lower than 70o
C (50 and 60o
C) for the feed
and 30o
C for the distillate were chosen for the following experiments.
14.4.2. Membrane scaling investigation
To further investigate the onset of rapid flux decline, a set of experiments was
performed in conjunction with stereomicroscope observation. Water flux and total
solids concentration as a function of time are shown in Fig. 45.

84
Fig.45, Water flux and total solids concentration as a function of time for experiment performed with GLS and
NaCl feed solutions. Labeled dashed lines correspond to stereomicroscope images of the feed side of the
membrane surface during the experiment. Feed and distillate temperatures were 50 and 30
o
C, respectively.
Feed volume was 1,5L, flow velocities were 0,8Lxmin
-1
and the PP22 membrane surface area was 89cm
2
. [7]
The dashed lines in Fig.45 indicate when pictures of scaling on the membrane
surface were taken. Images of the PP22 membrane surface were captured before
and after the onset of rapid flux decline. Similar to results shown in Fig. 43, the water
flux begins to rapidly decline at a total solids concentration of approximately
300,000mg/L.
Images of the different stages of scale formation on the feed side of the membrane
are also shown in Fig.5. The onset of membrane scaling is first visualized after 7.5h of
operating time, or when the feed solution concentration approached 250,000mg/L
total solids. Thereafter, crystals continued to precipitate on the membrane surface
and the water flux continued to decline. After approximately 9.5h of operation, or at
a bulk feed solution concentration of 300,000mg/L total solids, the membrane
surface was mostly covered with salt resembling NaCl crystals.
To further evaluate the effect of sparingly soluble salts on flux decline, results from
experiments performed with pure NaCl as the feed solution were super-imposed on
results from experiments performed with GSL water (Fig.45). Interestingly, similar to
results from experiments with GSL water, the same sharp decline in water flux
occurred during the experiments performed with the NaCl water, indicating that the
onset of homogeneous precipitation of salts, mainly NaCl, correlates to the onset of
rapid water flux decline.
Water flux was higher during experiments with NaCl feed than during experiments
with GSL water. Also, compared to the experiments with the GSL water, the rapid
decline in water flux is delayed in the experiments with NaCl feed. This can be
explained by further evaluating the complexity of the solution chemistry for the GSL
water. Sparingly soluble salts and organic matter are present in the GSL water (Table
22). OLI modeling results revealed that at a bulk GSL water feed solution

85
temperature of 50o
C, calcium species are the first to reach saturation (at
340,000mg/LTDS), followed by NaCl (at 400,000mg/L TDS). Curcioetal, found that
divalent calcium ions in the presence of humic acid form complexes with the
carboxyl functional groups and cause membrane scaling. The calcium scaling then
serves as nucleation sites for other species, such as NaCl. Therefore, scaling of
sparingly soluble salts and NaCl were both the cause of rapid water flux decline
during the experiments with GSL feedwater.
14.4.3. Extended scaling experiments
Water flux and distillate conductivity as a function of time and feed total solids
concentration are shown in Fig. 46a and b.
Fig.46, Water flux and distillate conductivity as a function of (a) elapsed time and (b) total solids concentration
in the feed with experiments with the TS22 the PP22 membranes using cartridge filtered GLS feed water. The
feed temperature was 50
o
C and the distillate streams was 30
o
C and the stream flow rates were 1,6 Lxmin
-1
. [7]
Long term batch experiments were performed in a unique operating mode to
evaluate the effect of membrane scaling over time. The first part of the cycle was
performed to evaluate how water flux decreases as concentration increases. Similar
to results obtained for the successive batch experiments in both sets of experiments,
the water flux gradually declined until the feed solution reached approximately
300,000mg/L total solids. The second part of the cycle (recirculation step) was
performed under constant conditions to evaluate membrane scaling over time and
its effects on water flux. During the recirculation step, water flux continued to
decline, further indicating that in addition to a reducing partial vapor pressure of the

86
feed solution, nucleation of sparingly soluble salts on the membrane also contribute
to the gradual decline in water flux.
14.5. Scaling mitigation techniques
Three unique operating techniques were investigated to mitigate membrane scaling.
These include reduced operating time interval, reduced operating time interval with
flow reversal, and reduced operating time interval with temperature reversal. The
PP22 membrane was chosen for these experiments because of its isotropic
structure. The operating temperatures were chosen because an accelerated
operating time and increased feed concentration can be achieved without wetting
the membrane and compromising its performance.
14.5.1. Mitigating rapid flux decline
The first technique to prevent scale formation during successive batch experiments
was to terminate the experiment before a rapid flux decline occurred. Water volume
recovered from the feed and distillate conductivity as a function of elapsed time are
shown in Fig. 47.
Fig.47, Water recovered and distillate conductivity as a function of elapsed time for the successive batch
experiments (runs) performed with the PP22 membrane and filtered GLS feed water. The numbers at the top
of each line represent the average water flux (Lm
-2
h
-1
) for each batch run. The experiments were conducted
with feed and distillate temperatures of 60 and 30
o
C, respectively. [7]
Each line represents a successive batch experiment, and the slope of each line
divided by the membrane area (0.0139m2
) is the average water flux (in Lm-2
h-1
)
during each batch experiment (labeled above each line). During the first five
successive batch experiments the water recovery only minimally changed. However,

87
during the 6th batch, scaling on the membrane and wetting of some pores caused a
rapid decline in water flux and a sharp increase in distillate conductivity. Following
these results, two new operating techniques were tested to mitigate and reverse
scaling of minerals on the membrane.
14.5.2. Flow reversal
Water volume recovered (i.e., distillate collected) and distillate conductivity as a
function of elapsed time are shown in Fig.48 and each line represents a successive
batch experiment.
experiments with alternating feed and distillate channels. The feed and distillate channels were alternated
three times each. S1 and S2 denote the initial feed and distillated sides, respectively whereas 1,2 and 3 denote
the first, second and third alternations of the feed and distillate sides. The numbers at the top of each line
represent the average water flux (Lm
-2
h
-1
) for each batch run experiment The experiments were performed
with the PP22 membrane and filtered GLS feed water and were conducted with feed and distillate
temperatures of 60 and 30
o
C, respectively at 1,6 Lxmin
-1
[7]
In this scale mitigation technique, the feed and distillate flow channels were
exchanged after each successive batch experiment. The average water flux (Lm-2
h-1
)
during each batch experiment is labeled above each line. The average water flux
during all experiments was 19.5 Lm-2
h-1
with a standard deviation of 1.44 Lm-2
h-1
,
indicating that membrane scaling was minimal. Following the first experiment, the
distillate conductivity increased. This increase in distillate conductivity was mostly
due to residual salts in the distillate hydraulic loop from the previous
cycle/experiment and/or dissolution of scalants that deposited on the membrane
surface and in the membrane pores. The cause for the different trends in distillate
conductivity on either side of the membrane is not well understood; however, it is
likely that the slight difference in surface characteristics on the opposite sides of the

88
membrane resulted in dissimilar scaling and wetting patterns on the membrane.
Additional research on membrane characteristics, nucleation kinetics, and scale
formation could provide further insight to this trend.
14.5.3. Temperature reversal
Water volume recovered and distillate conductivity as a function of elapsed time for
the third scale mitigation technique are shown in Fig.49.
experiments (runs) performed with the temperature reversal tecnique. The numbers at the top of each line
represent the average water flux (Lm
-2
h
-1
) for each batch run experiment The experiments were performed
with the PP22 membrane and filtered GLS feed water and were conducted with feed and distillate
temperatures of 60 and 30
o
C, respectively at 1,6 Lxmin
-1
[7]
The average water flux (in Lm-2
h-1
) during each batch experiment is labeled above
each line. In this technique, the temperature difference across the membrane was
reversed for a period of time before a new batch experiment was performed.
The average water flux during these experiments was 20.6 Lm-2
h-1
with a standard
deviation of 0.95 Lm-2
h-1
. The water flux slightly declined during the sixth experiment
and the distillate conductivity slowly increased. Overall, the water flux and salt
rejection were higher during this operating technique than the previous scale
mitigation techniques, and the use of freshwater to flush the feed channel was
eliminated.
Flow and temperature reversal techniques proved to be very effective in maintaining
water flux and mitigating membrane scaling. Compared to previous experiments
performed without scale mitigation techniques, experiments performed with the

89
flow and temperature reversal techniques resulted in sustained, high water fluxes
throughout batch concentration experiments. Also, these operating techniques were
performed without the use of chemicals (i.e., antiscalants, acids and bases) to
remove scalants, and without additional energy to cool the feedwater. Scale
mitigation via temperature reversal achieved the greatest average water fluxes,
rejected nearly 100% of non-volatiles for the first six batch concentration cycles and
was performed without additional water and energy inputs. Therefore, this scale
mitigation technique could prove to be very impactful in MD. A recent study by
Kesieme reported that addition of a 0.5 μm filter before the feed channel inlet was
effective in capturing precipitated salts that would have otherwise deposited on the
membrane. Hybridization of the proposed scale mitigation techniques with a 0.5 μm
filtration of the feed could further reduce membrane scaling and wetting in MD.
14.6. Efficiency of MD over natural evaporation
When considering replacement of evaporation ponds with DCMD, two central
considerations are the time and costs involved in concentrating brines. Production of
high-value minerals will continue to increase with growing demands, and acquisition
of land for additional ponds can be costly or in some cases impossible. Therefore, the
efficiency of natural evaporation of the Bear River Bay was compared to MD.
The net annual evaporation rate at the Bear River Bay is 1040 mm per year. From a
simple unit conversion, on average 2.85 mm of water is evaporated from the bay
each day. From results obtained in this research, DCMD can concentrate GSL water
at an average rate of 20 Lm-2
h-1
(Fig.49), and from a simple unit conversion,
approximately 480 mm of high-quality distillate water can be recovered from GSL
water each day using a DCMD membrane of the same area. Therefore, applying
DCMD to mineral production not only recovers high quality water, but also
accelerates the natural evaporation process in concentrating hypersaline solutions
by approximately 170 times.
In terms of land use, one acre (4047 m2
) of evaporation ponds could be replaced
with approximately 24 m2
of flat sheet DCMD membrane. Several studies have
shown that DCMD is an economically and environmentally competitive water
treatment process to RO when low-grade heat is utilized. A study by Al-Obaidani
estimated that operating DCMD with a heat recovery system could reduce water
cost to $0.64m-3
of water produced (≈40 kWh/m3
), making DCMD a competitive
membrane process to RO ($0.50m-3
of water produced). A more recent study
estimated that water production costs with DCMD could be further
reducedto$0.57m-3
when low-grade heat is utilized and carbon tax is applied.
Therefore, in addition to concentrating GSL water for mineral recovery, the high-
quality water produced can be sold to further off set operating costs.

90
14.7 Conclusions
DCMD was effective in concentrating GSL water to greater than 350,000mgL-1
.
Operating DCMD at high ΔT of 40o
C was not sustainable; the membrane
performance was compromised because of membrane scaling and pore wetting.
Consequently, operating DCMD in successive batch mode without the use of scale
mitigation techniques resulted in decreased membrane performance (i.e., lowered
salt rejection and water fluxes).
Flow reversal and temperature reversal are new operating techniques that proved
very effective in sustaining high water fluxes and membrane performance. The scale
mitigation techniques were effective in inhibiting homogeneous precipitation of salts
and disrupting nucleation of sparingly soluble salts on the membrane surface. Of the
three scale mitigation techniques, the temperature reversal technique was most
effective in maintaining high water fluxes(>420 Lm-2
h-1
) and high salt rejection. The
new techniques were simple to operate and very impactful in mitigating scaling.
Furthermore, the need for antiscalants and other chemicals used for membrane
cleaning was avoided.
Replacing natural evaporation ponds with DCMD can result in enhanced operations
and reduced environmental footprints. Operating DCMD with low-grade heat
recovered from the on-site chemical processing plant can drastically reduce MD
operating costs, and high-quality water recovered from the GSL water can aid in
offsetting operating costs.

91
15. References
1. Membrane distillation principle and applications (2011), Khayet & Matsuura,
Elsevier Publications
2. Membrane Distillation: Principle, Advances, Limitations and Future Prospects in
Food Industry, Onsekizoglu, www.intechopen.com
3. High recovery of concentrated RO brines using forward osmosis and membrane
distillation (2009), Riziero Martinetti, Childressa & Cath, Journal of Membrane
Science 331 31–39
4. Advances in Membrane Distillation for Water Desalination and Purification
Applications (2013), Camacho , Dumée, Zhang, Li, Duke, Gomez & Gray, Water
2013, 5, 94-196; doi:10.3390/w5010094
5. Desalination Using Membrane Distillation Experimental and Numerical Study
(2011), Alaa Kullab, Doctoral Thesis; Royal Institute of Technology SE-100 44
STOCKHOLM
6. Study on concentrating primary reverse osmosis retentate by direct contact
membrane distillation; D. Qu, J. Wang, B. Fan, Z. Luan & D. Hou; Desalination 247
(2009) 540–550
7. Sustainable operation of membrane distillation for enhancement of mineral
recovery from hypersaline solutions; K.Hickenbottom,T.Cath; Journal of Membrane
Science 454(2014) 426–435
8. Integration of accelerated precipitation softening with membrane distillation for
high-recovery desalination of primary reverse osmosis concentrate; D.Qua,
J.Wanga, L.Wangb, D.Houa, Z.Luana, B.Wangb; Separation and Purification
Technology 67 (2009) 21–25
9. Vacuum membrane distillation of seawater reverse osmosis brines; J.Mericq,
S.Laborie, C.Cabassud; Water research 44 (2010) 5260 - 5273

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