RO&MD booklet

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

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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.

4


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].

5


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.

11


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-1K-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.

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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/m2day. 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/(m2xh), which were lower

than the RO permeate fluxes (20-75 kg/(m2xh)). 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 38oC 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/(m2h). 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 = 60oC, Tp = 20

oC 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 = 60oC, Tp = 20

oC, 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 = 50oC, 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 20oC.

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 = 50oC 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 = 50oC 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 = 50oC, Re = 4500), (c&d) feed

temperature (Pp =500Pa, Re = 4500), and (e&f) Re number (Pp = 500Pa, Tf = 50oC). [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 70oC increases the permeate vapour

flux from 1.33 to 22.30 L h-1m-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 70oC, 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 = 50oC and Re = 4500).

60


Fig.33, Computed variation of permeate flux versus Knudsen permeability for different solutions

concentrations, (Pp = 500 Pa, Tf = 50oC 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 oC 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.

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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.

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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.

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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.

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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)),

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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

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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

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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.

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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) 30oC ,the feed temperature was 30-70

oC

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=40oC) and

low temperature differential (ΔT=20oC) were chosen for the successive batch

concentration experiments; the temperature of the distillate stream was kept at

30oC and the temperature of the feed stream was either 50 or70oC.

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 20oC is shown in

Fig.44.

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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 30oC and the feed

temperature was either 50oC (50/30) or 70

oC (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 20oC) and 20% (ΔT of 40oC) 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 20oC) and 38% (ΔT of 40oC)

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 20oC 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.

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At increased temperature differentials (ΔT of 40oC), a higher terminal GSL water

concentration was achieved with the PP22. Although experiments conducted with

the TS22 at a higher temperature difference (ΔT of 40oC) 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 30oC and the feed temperature was either 50

oC (50/30) or 70

oC (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 20oC

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 68oC 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 (<350oC); however, the potential

for wetting of PTFE membranes does increase with increasing feed temperatures

and salinities.

When operated at a higher ΔT of 40oC, 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 20oC, 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 70oC (50 and 60oC) for the feed

and 30oC 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.

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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 30oC, respectively.

Feed volume was 1,5L, flow velocities were 0,8Lxmin-1

and the PP22 membrane surface area was 89cm2. [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 50oC, 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 50oC and the distillate streams was 30

oC 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 30oC, 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-2h-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 30oC, 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-2h-1)

during each batch experiment is labeled above each line. The average water flux

during all experiments was 19.5 Lm-2h-1 with a standard deviation of 1.44 Lm-2h-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 30oC, respectively at 1,6 Lxmin

-1 [7]

The average water flux (in Lm-2h-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-2h-1 with a standard

deviation of 0.95 Lm-2h-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

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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-2h-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.

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14.7 Conclusions

DCMD was effective in concentrating GSL water to greater than 350,000mgL-1.

Operating DCMD at high ΔT of 40oC 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.

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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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