Potential of membrane distillation - a comprehensive review

Int. J. Water, Vol. 7, No. 4, 2013 317

Copyright © 2013 Inderscience Enterprises Ltd.

Potential of membrane distillation – a comprehensive review

Adnan AlHathal Al-Anezi* Centre for Osmosis Research and Applications (CORA), Chemical Engineering Department, Faculty of Engineering and Physical Sciences, University of Surrey, Guildford GU2 7XH, UK and Department of Chemical Engineering Technology, College of Technological Studies, The Public Authority for Applied Education and Training (PAAET), P.O. Box 42325, Shuwaikh 70654, Kuwait E-mail: [email protected] *Corresponding author

Adel O. Sharif and Mohammed I. Sanduk Centre for Osmosis Research and Applications (CORA), Chemical Engineering Department, Faculty of Engineering and Physical Sciences, University of Surrey, Guildford GU2 7XH, UK E-mail: [email protected] E-mail: [email protected]

Abdul Rahman Khan Department of Environment Technology and Management, College for Women, Kuwait University, P.O. Box 5969, Safat 13060, Kuwait E-mail: [email protected]

Abstract: Membrane distillation (MD) is a recent and unique separation technology, in use in the process industry. The process of separation in MD involves the simultaneous heat and mass transfer through a hydrophobic semi permeable membrane, using thermal energy. Consequently a separation of the feed solution into two components – the permeate or product and the retentate or the return stream occurs. MD utilises low grade or alternative energy, e.g., solar energy, geothermal energy, etc., as a source and is the most cost effective separation technology. Hence the process has come to acquire the attention and interest of researchers, experimentalists and theoreticians all over the world. This article is a comprehensive review of the prominent research in the field of MD technology, including its basic principle, MD configurations, area of applications, membrane characteristics and modules, experimental studies involving the effect of main operating parameters, MD energy and economic, fouling and long-term performance.

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Keywords: membrane distillation; MD; MD configurations; membrane modules; parameters effecting MD; MD energy and economic; fouling and long-term performance.

Reference to this paper should be made as follows: Al-Anezi, A.A., Sharif, A.O., Sanduk, M.I. and Khan, A.R. (2013) ‘Potential of membrane distillation – a comprehensive review’, Int. J. Water, Vol. 7, No. 4, pp.317–346.

Biographical notes: Adnan AlHathal Al-Anezi is a faculty member at Department of Chemical Engineering Technology, Public Authority for Applied Education and Training (PAAET) at Kuwait, and post-graduate student at Chemical and Process Engineering Department, University of Surrey Guildford, UK.

Adel O. Sharif is a Founder Director of the Centre for Osmosis Research and Applications, (CORA) at the Faculty of Engineering and Physical Sciences, University of Surrey, UK and is the winner of the Royal Society 2005 Brian Mercer Senior Award for Innovations in Science and Technology.

Mohammed I. Sanduk is a member of the Centre for Osmosis Research and Applications, (CORA) at the Faculty of Engineering and Physical Sciences, University of Surrey, UK. His current researches within CORA lie in desalination, electrostatic ions separation (CDI), direct use of solar and renewable energies in desalination and water treatment.

Abdul Rahman Khan is a Senior Researcher at Kuwait Institute for Scientific Research for five years, and a faculty member at Chemical Engineering Department Kuwait University for 12 years.

1 Introduction

Desalination technology has been a source of pure water for centuries. Researches in the field have been constantly creating more energy efficient and cost effective methods. Among the recent technologies, membrane distillation (MD) has the advantage of performing at moderate temperatures and pressure (Mannella et al., 2010). The MD is a thermally driven separation process and it is economical in terms of energy, since the heat source for the process can be solar energy and also energy is recovered continuously (Mannella et al., 2010; Li et al., 2008). During the process of MD, a hot saline is brought in contact with a hydrophobic membrane, which selectively allows water vapour to diffuse through, restricting the flow of liquid and hence, dissolved salts through its pores. The mass transfer of water vapour through membrane pores is driven by the vapour pressure difference, as well as the temperature difference between the two sides of the hydrophobic membrane, i.e., the feed side and the permeate side (Mannella et al., 201; Al-Obaidani et al., 2008). MD is innovative because it brings together the key components of conventional distillation and other membrane separation processes such as multistage flash (MSF) and reverse osmosis (RO). Besides being an economical, energy efficient and eco-friendly process, several other characteristics of MD make it promising, such as

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1 compact process size

2 theoretically, it promises 100% rejection of ions, macromolecules, colloids, biological cells and other non-volatile impurities

3 operating temperatures considerably below boiling (varies in the range of 30°C–90°C) allows the process to take advantage of low-grade waste heat, unlike conventional distillation

4 the membrane and the process solution interact less

5 MD is less demanding of the membrane’s mechanical strength

6 lower operating pressure than pressure-driven processes (atmospheric to a few hundreds of kPa)

7 alternative energy sources such as solar, wave or geothermal could be used to drive the process

8 MD process suffers less fouling due to its relatively large pore size required compared to the conventional membrane separation processes such as RO (Curcio and Drioli, 2005; Smolders and Franken, 1989; Lawson and Lloyd, 1997; Couffin et al., 1998; de Andrés et al., 1998).

Figure 1 A general scheme of the MD process

Source: Curcio and Drioli (2005)

MD can take on four configurations namely direct contact, air gap, sweep gas, vacuum MD, as seen in Figure 1, and the difference between these configurations depends on the way in which the vapour is condensed and/or removed from the membrane module, where all four configurations have advantages and disadvantages, depending on their applications with the feed solution to be treated (Curcio and Drioli, 2005; Lawson and Lloyd, 1997). Due to the above advantages MD process has been applied successfully in the field of seawater desalination (Basini et al., 1987; Martinez and Florido, 2001; Peng et al., 2005), removal of heavy metals (Zolotarev et al., 1994), acid saline effluent (Gryta et al., 2006) and ammonia (Xie et al., 2009) from waste water, recovery of dilute

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acetone-butanol-ethanol (ABE) solvents (Banat and Al-Shannag, 2000) and separation of ethylene glycol (EG) (Mohammadi and Akbarabadi, 2005) from aqueous solutions. Moreover, it has been used effectively in biomedical and pharmaceutical applications (Sakai et al., 1988; Ding et al., 2008; Zhao et al., 2008). The nuclear industry also employs MD in nuclear desalination (Khayet et al., 2006a) and concentration of radioactive solutions (Zakrzewska-Trznadel et al., 1999), purification and concentration of chemicals (Tomaszewka, 2000), food processing among others like fruit juice concentration (Bandini and Sarti, 2002).

2 Principle of MD

The basic concept of MD lies in the traditional distillation process itself, both being based on vapour-liquid equilibrium, and involves a phase change in the process. However, the differentiating criterion is the temperature of operation, which in MD is much lower than the boiling point of the solution being separated. The criteria essentially describing a process as MD are (Smolders and Franken, 1989; Lawson and Lloyd, 1997; Bryk and Nigmatullin, 1994):

1 the membrane should be highly hydrophobic and highly porous

2 membrane not wetted by process liquids

3 no capillary condensation inside the membrane pores

4 exclusive transport of vapours through membrane pores

5 vapour-liquid equilibrium of process liquid components should not be altered by the membrane

6 direct contact of at least one side of the membrane with process liquids

7 driving force for each component is the partial pressure gradient in the vapour phase.

3 MD configurations

MD can take four configurations, and the difference between these configurations depends on the way in which the vapour is condensed and/or removed from the module.

3.1 Direct contact membrane distillation

A direct contact membrane distillation (DCMD) process is characterised by a micro-porous membrane separating the two chambers of MD, one side having the feed water, while the other side has permeate and the cooling liquid, both in direct contact with the membrane. There is a liquid vapour interface on either side of the membrane, and only vapour from the feed side passes through the membrane to the permeate side, where it condenses to form the output. A higher temperature on the feed water side compared to permeate side is the driving force, which enables the vapour to cross the membrane. Membrane being the sole barrier in this case; direct contact means that both the feed water and the permeate are in direct contact with the membrane in the

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chambers. High permeate flux from 1 to 10 (kg/m2 h) and good salt rejection have been achieved for DCMD system for salt solutions in the range of 35,000 and 100,000 ppm (Summers et al., 2012). Unfortunately, the main drawback of this process is the high heat lost by conduction through the membrane (Curcio and Drioli, 2005; Lawson and Lloyd, 1997; Alkhudhiri et al., 2012a). The mechanism of DCMD is illustrated in Figure 2.

Figure 2 Direct contact membrane distillation (see online version for colours)

3.2 Air gap membrane distillation

The mechanism of air gap membrane distillation (AGMD) is shown in Figure 3. AGMD configuration was first proposed by Carlsson (1983), and as the name suggests is characterised by the presence of an air gap, on one side of the membrane. In AGMD, only the feed solution is in direct contact with the membrane, while on the permeate side, an air gap is introduced, separating the cooling chamber from the membrane. Thus in AGMD, the permeate is condensed on a cooler surface, compared to DCMD, which also alleviates the loss of heat energy by conduction through the membrane. The vapour molecules crossing to the permeate side of membrane emerge out of the membrane pores, travel across the air gap, and condense on the plate. The condensed liquid is drained out of the base of the air gap by gravity. The air gap provides an additional resistance to mass transfer in AGMD, which on one hand is the major drawback of the configuration. On the other hand, it enables the process to be also used for the separation of volatile compounds, such as alcohol from aqueous solution. The wetting probability is much lower in AGMD also due to the absence of direct contact between membrane and permeate side due to presence of the air gap. Hence, the risk of wetting on the permeate side is much reduced. The width of air gap being much more than that of membrane, added to the lower heat conductivity of air, then the passage of vapour across the air gap becomes the determining step of the process. AGMD has a capability of recovering the latent heat of vaporisation, and has much higher thermal energy efficiency compared to DCMD (Lawson and Lloyd, 1997; Tomaszewska, 2000; Wirth and Cabassud, 2002). Moreover, AGMD has a slight higher flux ranging from 10 (L/m2 h) up to 65 (L/m2 h) for inlet feed temperature near to 70ºC (Summers et al., 2012).

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Figure 3 Air gap membrane distillation (see online version for colours)

3.3 Sweep gas membrane distillation

Sweep gas membrane distillation (SGMD), also referred to as air stripping MD, is based on the mechanism (Figure 4) of removal of permeate by a flow of sweeping gas and its subsequent condensation done externally. SGMD was first employed for water desalination by Basini et al. (1987). In AGMD, air is blown into the cooling chamber and onto the membrane surface tangentially, instead of keeping a still air cooling air gap between the membrane and cooling chambers, as in AGMD. Like AGMD, the flowing air removes volatile impurities from the solution (Findley, 1967), it further has the advantage of offering least resistance to the passage of vapour through it. However, the sweeping gas dilutes the permeate vapour, and this leads to more stress on the condenser, both in terms of capacity and in energy consumption, which is a major limitation of SGMD (Alklaibi and Lior, 2004; Soni et al., 2008).

Figure 4 Sweep gas membrane distillation (SGMD) (see online version for colours)

3.4 Vacuum membrane distillation

Vacuum membrane distillation (VMD) is an improvisation on SGMD, wherein instead of the mixture of sweeping gas and the permeate vapour, the vapour itself is removed alone by vacuum pump, and is subsequently condensed externally. The mechanism of the VMD

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is shown in Figure 5. The mechanism makes the method appropriate for separation of various volatile compounds from their aqueous solutions or a mixture of the same (Alklaibi and Lior, 2004; Banat et al., 2003). It is only recently that it was applied for seawater desalination and treatment of RO brines (Martínez-Díez et al., 1998; Cerneaux et al., 2009; Schneider and van Gassel, 1984; Xu et al., 2006). Here, the feed solution is maintained at a pressure lower than the minimum entry pressure (LEP), when in contact with membrane. On the other side, the permeate is also maintained at a lower pressure than the equilibrium vapour pressure using a vacuum pump. This enables a VMD process to achieve the best performance in terms of flux compared to other MD configurations. Also, since the permeate pressure is kept low, the heat transfer through the membrane is negligible, and therefore VMD is highly energy efficient (Wang et al., 2009). For example, a novel VMD system for desalination achieved a high permeate flux of 71 (kg/m2 h) operating at a hot feed temperature of 85ºC and under vacuum of 60–66 cm Hg (Diban et al., 2009), while another VMD system for water desalination tested using two types of ceramic membranes and operating at 40ºC feed temperature and permeate pressure of 300 Pa was able to achieve a higher flux of 146 and 180 (L/day m2) (Fujii et al., 1992).

Figure 5 Vacuum membrane distillation (see online version for colours)

4 Membranes characteristics

Besides all the advantages of MD, special membranes for this process have not been provided. Membrane characteristics are the major determinants of a MD process, and involve the use of hydrophobic (non-wetting), micro-filtration membranes exclusively, because most of the specifications required by MD processes are available from those membranes (El-Bourawi et al., 2006; Zhigang et al., 2005). These membranes are mainly fabricated from chemical resistance polymers, such as polyethylene (PE), polytetrafluoroethylene (PTFE), polypropylene (PP) and polyvinylidene fluoride (PVDF) (Lawson and Lloyd, 1997; El-Bourawi et al., 2006; Zhigang et al., 2005). Table 1 shows some commercial membranes commonly used in MD with their characteristics. In general, the membrane used in the MD process should have higher permeability, i.e., low resistance to mass transfer, lower membrane thickness, high liquid entry and low thermal conductivity to prevent heat loss across the membrane. In addition, the membrane should have good physical and thermal stability in extreme temperatures, and exhibit excellent

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chemical resistance, such as acids and bases (Alkhudhiri et al., 2012a; Zhigang et al., 2005; Dow et al., 2008). The membranes used in the MD system is generally characterised by four parameters, i.e., the thickness δm (m), the mean pore size, diameter dp or radius rp (m), the porosity ε (defined as the volume of the pores divided by the total membrane volume) and the tortuosity τ (defined as the ratio of pore length to membrane thickness), where these four parameters strongly affect the membrane permeability (Smolders and Franken, 1989; Zhigang et al., 2005). Membranes with average pore diameter in the range of 0.1 µm to 1 µm (El-Bourawi et al., 2006; Scheer et al., 1999; Khayet et al., 2008) can be used for MD, in order to allow the mass flow of vapour, the pores must be large, yet small enough to inhibit the intrusion of water under operation (Lawson and Lloyd, 1997; El-Bourawi et al., 2006). Table 2 shows the usual membrane characteristic values used in MD (Zhigang et al., 2005). Table 1 Some commercials membranes commonly used in MD with its characteristics

Membrane materials

Porosity (ε) (%)

Mean pore size (μm)

Thickness (δm) (mm) Reference

PTFE 70 0.2 70

PTFE 90 0.2 64 Phattaranawik et al. (2003)

PVDF 89 0.45 77

PTFE 60 0.2 60 Gostoli and Sarti (1989)

PVDF 75 0.45 110 Schofield et al. (1990)

PTFE 80 0.45 60 Martínez-Díez et al. (1998)

PP 80 0.1, 0.2, 0.45 60 Martıcnez et al. (2002)

PTFE TF-200 0.2

PTFE TF-450 80 0.45 178 Rincón et al. (1999)

PTFE TF-1000 1.0

PTFE 85 0.1 150 Zhu (1999)

PVDF 80 0.2 125 Phattaranawik et al. (2001)

PP 40 0.03 31 Wang et al. (2001)

PP 0.22 150

PTFE 70 0.22 Cath et al. (2004)

PTFE 0.45 175

PVDF 1.0

PTFE Teflon-200 80 0.2 60 Rodríguez-Maroto and Martínez (2005)

PP 70 0.2 150 Narayan et al. (2002)

PVDF 75 0.45 110 Banat and Simandl (2000)

PE 66.3 0.087 50

PP 53.3 0.074 50 Li et al. (2003)

PP 50.0 0.044 65

PP 47.3 0.056 42

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Table 2 The usual range of membrane characteristics values used in MD

Mean pore diameter Porosity Thickness Tortuosity

0.1–1.0 µm 40%–90% 20–200 µm 1.5–2.5

5 Membrane modules

MD processes have used a large variety of membrane modules, including spiral wound and plate and frame modules (mainly using flat-sheet membranes), tubular and hollow fibre modules (mainly using tubular membranes) (Curcio and Drioli, 2005; El-Bourawi et al., 2006).

5.1 Spiral wound membrane module

A spiral wound module (Figure 6) is characterised by feed flow channel spacer, the membrane and porous support, where all of them are enveloped and wound around a perforated centre. The direction of the feed solution is axial across the membrane surface, while that of permeate is radial and exits towards the central tube. The basic principle of this type of module is to pack a large area into a small volume. Gore proposed a spiral wind module as early as 1982. Bier and Plantikow, 1995, Koschikowski et al. (2003) and Winter et al. (2011) have used a PTFE membrane spiral wound module for a solar powered desalination system.

Figure 6 Spiral wound membrane module (see online version for colours)

5.2 Plate and frame membrane module

In this type, the module uses a set of two plates placed at the ends, and the entire set of equipment, e.g., porous support plates, and spacers are stacked within this casing. Laboratory scale MD modules usually use a flat sheet membrane setup (Figure 7), since these are more flexible than tubular or hollow fibre setups. It is also easy to clean, remove, examine, and replace. The plate and frame module of flat sheet membranes has been used widely by many researchers for desalination and water treatment purposes (Lawson and Lloyd, 1996; Srisurichan et al., 2006; Andersson et al., 1985; Kimura et al.,

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1987; Khayet et al., 2004a; Srisurichan et al., 2005; Termpiyakul et al., 2005; El-Abbassi et al., 2009).

Figure 7 Plate and frame module for flat sheet membrane

5.3 Tubular membrane module

The tubular membrane module, as shown in Figure 8, is made up of several tubular membranes arranged as tubes. The tubular membrane module is more attractive than other modules, because of a high membrane surface area; therefore, in MD applications, it is suitable for fluids of high viscosity, and has the added advantages of permitting high flow rates along with reduced fouling, easy cleaning, and low tendency to polarisation phenomena. However, the main drawback of this type of module is low packing density, which results in a high cost per module. Schofield et al. (1987) utilised the tubular module in DCMD for water desalination. They concluded that the module was shown potentially to be the most efficient. Cerneaux et al. (2009) used tubular ceramic membranes for treating NaCl aqueous solution in DCMD, AGMD and VMD configurations.

Figure 8 Tubular membrane module (see online version for colours)

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5.4 Hollow fibre module

Hollow fibre modules (Figure 9) have thousands of hollow fibres installed and sealed inside a shell and tube. The hollow fibre modules can be employed in MD because of very high packing density and high specific surface area, which improves the vapour diffusion across the membrane. On the contrary, the module has very high chances of fouling, and maintenance and cleaning is difficult. A hollow fibre module was developed by ENKA for MD, as early as 1984 (Schneider and van Gassel, 1984). Li et al. (2003), Xu et al. (2006) and Wang et al. (2009) have used a hollow fibre module for water desalination. Moreover, experimental studies have been performed using the hollow fibre module for recovering the mean aroma compound (Diban et al., 2009) and for concentrating apple juice and alcohol respectively (Fujii et al., 1992; Laganà et al., 2000).

Figure 9 Hollow fibre membrane module (see online version for colours)

6 Experimental studies on MD

The fundamental operating parameters of the MD processes listed below are the most parameters that influence the permeate flux in the MD system.

6.1 Effect of feed temperature

Feed temperature is an extensively studied MD parameter in varied MD configurations, and has a remarkable effect on the permeate flux, as can be seen in Table 3. The influence of feed temperatures was investigated for a wide range of 17.5ºC to 90ºC, provided that the upper limit remains well below the boiling point of the feed. From the Antoine equation, the vapour pressure increases exponentially with temperature. Therefore, the general effect observed was an exponential increase in MD flux with rising feed temperatures (Banat and Al-Shannag, 2000; Alklaibi and Lior, 2004; Banat et al., 2004). The obvious reason for this enhancement was the increase in the driving force of the MD process, i.e., rise in trans-membrane vapour pressure gradient. The same can be lowered by reduced membrane selectivity (Alklaibi and Lior, 2004; Hsu et al., 2002). Several studies concluded that operating the MD process under high feed temperatures increases the evaporation efficiency (Alklaibi and Lior, 2004; El-Bourawi et al., 2006;

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Izquierdo-Gil et al., 1999). However, high feed temperatures, i.e., 90ºC, may increase the risk of scaling problems, reduces the membrane selectivity, and increases the temperature polarisation effect (El-Bourawi et al., 2006; Gryta and Karakulski, 1999). Table 3 The effect of feed temperature (Tf) on permeate flux

MD configuration

Feed solution

Feed flow

(L/min) Tf(cC)

Permeate flux

≈ (kg/m2h) Tp(°C) Ref.

VMD Pure water

1 gpm 30–75 1–9 (mol/m2s)

20 Lawson and Lloyd (1996)

DCMD NaCl (3%)

3.3 5–45 1–40 20 Hsu et al. (2002)

AGMD NaCl (3%)

3.3 5–45 0.5–8

DCMD Salt (NaCl)

2.0 61–81 18–42 21 Schofield et al. (1990)

Sugar 22–55 DCMD Brine 3.3

cm/s 40–60 0.25–2.25 30 Ohta et al. (1990)

AGMD Seawater 5.5 40–70 1–7 20 Banat and Simandl (1998)

SGMD Pure water

0.15 m/s

40–70 3.6–18 20 Khayet et al. (2000a)

AGMD Tap water 3.8 35–75 3–23 20 Liu et al. (1998) DCMD Pure

water 20

cm3/s 20–50 2.9–25.2 14 Martinez-Diez and

Florido-Diaz (2001) DCMD Orange

Juice 20

kg/min 25–45 3–11 20 Calabro et al. (1994)

LGMD Water supply

0.33 40–70 3.2–10.5 10 Ugrozov et al. (2003)

DCMD Pure water

16 cm3/s

17.5–31

3.6–32.4 14 Martínez-Díez et al. (1999)

DCMD NaCl (5 wt%)

0.6 30–70 10–80 20 Adnan et al. (2012)

VMD NaCl (300 g/L)

1.8 25–55 5.5–11 20 Safavi and Mohammadi (2009)

AGMD HNO3 (4 M)

0.05 60–90 0.4–3.7 15 Matheswaran et al. (2007)

6.2 Effect of feed flow rate

In most cases, increasing the feed flow rate leads to higher permeate flux, as can be shown in Table 4. Higher permeate flux results, because of two main reasons; first, at a given temperature, the Reynolds number increased with increasing feed flow rate, and second, since higher flow velocity results in minimising the boundary layer resistance and consequently maximising the convective heat transfer coefficient and the temperature polarisation coefficient to increase. As a result, higher permeate flux can be achieved

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(El-Bourawi et al., 2006; Srisurichan et al., 2006; Li and Sirkar, 2005; Phattaranawik and Jiraratananon, 2011). Alklaibi and Lior (2004) concluded that increasing the feed flow rate threefold leads to increasing the permeate flux by about 1.3 times, which is less than half the effect of feed temperature. Izquierdo-Gil et al. (1999) indicated that the temperature and concentration polarisation decreased when increasing the feed flow rate. Moreover, increasing the feed flow rate promotes turbulence, which consequently increases the heat transfer coefficients and brings the membrane surface temperature closer to the corresponding bulk temperature (Khayet et al., 2010). Most studies mentioned that the permeate flux increases to an asymptotic value with increasing the feed flow rate until it approaches a certain limit, where the effect of the feed flow rate does not change the permeate flux any more (Alklaibi and Lior, 2004; Sudoh et al., 1997). Table 4 The effect of feed flow rate on permeate flux

MD configuration

Feed solution

Flow rate(L/min) Tf(ºC)

Permeate flux ≈

(kg/m2 h) Ref.

DCMD NaCl (25 wt %)

1–2 81 32–42 Schofield et al. (1990)

Sugar (30 wt %)

38–55

AGMD Seawater 2–5 60 2.5–3.3 Banat and Simandl (1998)

SGMD NaCl (1 M) 0.12–0.2 m/s

65 19.0 Khayet et al. (2000a)

DCMD Juice (108 g/L)

2–5 kg/min 45 5.76–9.7 Calabro et al. (1994)

VMD NaCl (100 g/L)

0.9–1.8 55 12.2–14.4 Safavi and Mohammadi (2009)

AGMD HNO3 (4 M)

0.05–0.2 80 2–5.5 (L/m2h)

Matheswaran et al. (2007)

VMD Acetone (5 wt%)

0.15–2.6 35 12.6–21.6 Bandini and Sarti (1999)

VMD TCA (620 ppm)

0.1–1 L/h 50 2.5–45 (g/m2h)

Wu et al. (2006)

DCMD Sucrose 45 wt%

2.2–3 m/s 36 7.9–10.8 Martínez and Rodríguez-Maroto

(2007)

DCMD Arsenic (394 ppb)

0.028–0.062 (m/s)

60 29.16–46.24

Pal and Manna (2010)

DCMD Orange Juice 11.5º Brix

0.5–1.2 50 3.6–6.5 Deshmukh and Tajane (2010)

DCMD NaCl 5,000 ppm

1–4 70 2.5–5.5 Al-Hathal Al-Anezi et al. (2012)

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6.3 Effect of feed concentration

As can be seen in Table 5, the effect of feed concentration on the permeate flux depends on the MD configuration used. It is established that MD can be used with highly concentrated solutions, even with non-volatile solutes and a supersaturated solutions; unlike other membrane processes, especially the pressure driven methods, there is no considerable drop in membrane permeability (Lawson and Lloyd, 1997; El-Bourawi et al., 2006; He et al., 2011; Banat and Simandl, 1994). On the other hand, Yun et al. (2006) investigated the effect of high concentration NaCl solution (24.68 wt.%) and found a remarkable variation in permeate flux with time. The flux reduction in concentrated solution is attributed to the fact that the higher the concentration of the NaCl solution, the higher the boiling point. Consequently, the high salt concentration, and increasing temperature and concentration polarisation at membrane surface, reduces the vapour pressure gradient, which is the driving force of the MD process (He et al., 2011; Banat and Simandl, 1994; Martıcnez-Dıcez and Vázquez-González, 1999; Qtaishat et al., 2008). The effect of NaCl solution concentration on permeate flux was investigated by He et al. (2011), who found that the permeate flux decreased slightly from 17.8 L/m2 h to 14.2 L/m2 h, when the concentration of NaCl increased from 1 to 10 wt.%. Moreover, Qtaishat et al. found that about 12% permeate flux reduction occurred when NaCl concentration increased from 0 to 2 M concentration (Qtaishat et al., 2008). The increase in the feed concentration leads to increase in the temperature polarisation and the formation of a boundary layer on the feed side of the membrane surface, due to increased concentration polarisation, which also contributes to lowering the vapour pressure and reducing the performance of the MD process (Lawson and Lloyd, 1997; Safavi and Mohammadi, 2009). Schofield et al. (1990) studied the influence of 25 wt.% of NaCl solution and 30 wt. % of sucrose solution. Under the same conditions of feed velocity, temperature and same membrane type, they found that the flux reduction of sucrose solution is less than that for salt solution. This is mainly attributed to the higher sucrose molecular weight fraction, resulting in less reduction of vapour pressure than for NaCl, and the increase in viscosity plays an important factor in the flux reduction for sucrose solution. The effect of the high concentration of four salts, i.e., NaCl, MgCl2, Na2CO3 and Na2SO4 on permeate flux was investigated by Alkhudhiri et al. (2012b). They found that the flux reduction can be referred to the vapour pressure reduction across the membrane. Martinez (2004) concluded that the reduction in the permeate flux is mainly due to the reduction in water activity, which decreases as the solute concentration increases. However aqueous solutions of volatile substances, such as alcohols (methanol, ethanol and isopropanol), have an entirely different influence on MD flux.

The specific effect being a function of thermodynamic characteristics of the volatile solute and the nature of its interactions with water. Generally, higher concentration of a volatile solute leads to higher trans-membrane partial pressure of the compound due to its increased concentration on feed side, and as a consequence isopropanol has the highest vapour pressure. This results in higher permeate flux for isopropanol solution (Garcıca-Payo et al., 2000).

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Table 5 The effect of feed concentration on permeate flux

MD type Tp(ºC) Concentration

Flow rate

(L/min) Tf(ºC)

Permeate flux

≈ (kg/m2 h) Ref.

DCMD NaCl 0–9.5 2 81 62–43 Schofield et al. (1990) Sugar 0–4.1 mol % 70–55

DCMD NaCl 0–2 mol/L 16 cm3/s 31 32.4–24.8 Martínez-Díez et al. (1999)

VMD NaCl 100–300 g/L 1.8 55 13–9 Safavi and Mohammadi (2009)

AGMD HNO3 0.5–9 (M) 0.05 80 5–1.5 Matheswaran et al. (2007)

DCMD Arsenic 0–1200 ppb 0.028 m/s

60 27.5–23.6 Pal and Manna (2010)

DCMD Juice 6–225º Brix 1.2 50 8.6–4.3 Deshmukh and Tajane (2010)

DCMD NaCl 0–5000 ppm 4 60 3–7 Al-Hathal Al-Anezi et al. (2012)

CH4 30–200 3.9–4.6 AGMD C2H6O 30–150 2 35 4–4.8 Garcıca-Payo et al.

(2000) C3H8O 30–100 g/L 4.15–5 VMD Brine 50–3,000 g/L Re =

4,500 50 45–28 Martlnez (2004)

6.4 Effect of coolant permeate temperature

It was established that for a constant feed temperature, the permeate flux increases by decreasing the temperature at the permeate side, which is related to the rise in trans-membrane vapour pressure (Calabro et al., 1994). Moreover, the decrease in the permeate temperature is almost double (Gunko et al., 2006; Jönsson et al., 1985; Kurokawa and Sawa, 1996), and more than double the permeate flux (Alklaibi and Lior, 2004). On the contrary, Pangarkar and Sane (2011), Matheswaran et al. (2007) and Banat and Simond (1998), found that at fixed feed temperature, the changes in the permeate temperature have small effect on permeate flux and can be neglected. This result can be referred to the fact of low vapour pressure variation at low temperatures. The nature of the effect of permeate temperature on efficacy of MD is a function of MD configuration used, as can be seen in Table 6. Therefore, in DCMD, a decrease in permeate temperature is a consequence of higher permeate flux (Lawson and Lloyd, 1997; Alklaibi and Lior, 2004; Calabro et al., 1994; Gunko et al., 2006; Jönsson et al., 1985; Kurokawa and Sawa, 1996). However, the same does not hold true for AGMD, where the dominating factor governing heat transfer is the heat transfer coefficient, and therefore the effect of permeate temperatures on flux are minimal and can be neglected (Banat and Simandl, 1994; Banat and Simandl, 1998; Matheswaran et al., 2007; Pangarkar and Sane, 2011). Several researchers mentioned that to obtain higher MD flux, a more convenient option is

332 A.A. Al-Anezi et al.

to increase the temperature of feed side rather than decrease that of permeate side (Lawson and Lloyd, 1997; Alklaibi and Lior, 2004; Gunko et al., 2006). Table 6 The effect of coolant permeate temperature (Tp) on permeate flux

MD type Tp(ºC) Concentration Flow rate

(L/min) Tf(ºC) Permeate

flux ≈ (kg/m2 h)

Ref.

AGMD 7–30 35,000 ppm 5.5 60 45 3.9–3 Banat and Simandl (1998)

DCMD 25 108 g/L Juice 2–5 kg/min

5.4–9 Calabro et al. (1994)

20 6.1–9.7 AGMD 10–25 HNO3 (4 M) 0.05 80 2–2.4 Matheswaran et al.

(2007) DCMD 30–20 Apple Juice 0.8 m/s 60 21–19 Gunko et al. (2006) 30–10 21–17 AGMD 10–25 4 g/L 55 L/h 60 22.5–21 Pangarkar and Sane

(2011) 35 g/L 12.5–11

6.5 Effect of temperature difference and mean temperature

Trans-membrane vapour pressure is the driving force of all configurations of the MD process. This can be maintained in DCMD, AGMD and SGMD by the temperature difference maintained on the feed side and permeate side of the membrane, while the driving force of (VMD) is maintained by applying a vacuum at the permeate side (Lawson and Lloyd, 1997; Alklaibi and Lior, 2004). Under conditions of constant mean temperature, permeate flux increases linearly with increase in interfacial temperature difference (Lawson and Lloyd, 1997; Schofield et al., 1987; Pangarkar and Sane, 2011; Nene et al., 2002; Kim et al., 2004; Velazquez and Mengual, 1995). This result can be attributed to the increase of equilibrium vapour pressure, and thereby the increases in trans-membrane driving force (Kim et al., 2004). Kubota et al. (1988) found that when the temperature difference increases, the conductive heat loss per unit permeate flux decreases. However, the heat efficiency increases. Chen et al. (2009) concluded that increasing the interfacial temperature gradient will increase the diffusion coefficient, and consequently, increasing the permeate flux. Conversely, under conditions of constant temperature difference, the permeate flux increases exponentially with rise in mean temperatures. This fact is attributed to the exponential relationship in the Antoine equation between mean temperature and vapour pressure difference (the driving force) (Alklaibi and Lior, 2004; Banat et al., 2004; Deshmukh et al., 2011a).

6.6 Effect of permeate flow velocity

Permeate velocity effects can be analysed exclusively, only for the configuration of DCMD and SGMD (Ding et al., 2006). In the presence of volatile components on the permeate side of the membrane module, there is increased heat transfer by reduction of temperature and concentration polarisation effects, result from higher permeate

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velocities. With the rise in heat transfer coefficient on the permeate side, the membrane surface temperature reaches a value closer to permeate side bulk temperatures, and thus the driving force and consequently permeate flux increases (Khayet et al., 2000a, 2000b, 2002, 2003; Zuo et al., 2011; Yang et al., 2011). Yu et al. (2011) pointed out that the thickness of the thermal boundary layer is reduced by increasing feed and permeate velocity, which consequently, enhances the transmembrane flux. However, increasing the permeate velocity will decrease the thermal efficiency due to the heat loss on the permeate side. On the other hand, Banat and Simandl (1994, 1998), Khayet et al. (2004b) and Deshmukh et al. (2011b) concluded that the effect of permeate velocity on the permeate flux is much lower than that of the feed velocity. Therefore, the concentration polarisation can be neglected at permeate side. In DCMD, the permeate fluid is usually distilled water (liquid). Hence, permeate velocity effects are prominent here, compared to SGMD, where the permeate fluid is air (gases). This is precisely due to the distinct physico-chemical properties, and specifically, the thermal conductivities of the two forms of matter. In SGMD, in the presence of non-volatile components, temperature polarisation is confined to the permeate side, while the flow velocity increase leads to a rise in flux till an optimum value is achieved. Further increase in velocity causes a decline in flux. Thus, in DCMD and SGMD, the value of permeate flow velocity should be carefully optimised, so as to obtain maximum efficiency (El-Bourawi et al., 2006; Alklaibi and Lior, 2004; Zuo et al., 2011; Bui et al., 2010).

7 Energy in MD

7.1 Thermal energy efficiency

MD is a thermal process; therefore, heat economy is an important and crucial issue. MD technology is still at its infancy, in terms of thermal energy and cost estimation. Most MD researchers have focused on maximising the performance of MD in terms of membrane flux rather than minimising the energy consumption and cost. From the standpoint of energy expenditure, the MD process is characterised by thermal energy efficiency, which is defined as the ratio of latent heat of vaporisation to the overall heat (latent and conduction) introduced to the membrane module (Al-Obaidani et al., 2008; Smolders and Franken, 1989; Qtaishat et al., 2008; Calabro et al., 1991). Hogan et al. (1991) and Calabro et al. (1991) pointed out that the thermal energy efficiency can be improved by increasing the driving force on the membrane. This improvement can be achieved by reducing the permeate temperature at constant feed temperature, where this method has raised the efficiency from 8% to 14%. On the other hand, Alklaibi and Lior (Alklaibi and Lior, 2005) stated that an increase in the hot solution feed temperature from 40ºC to 80ºC enhances the thermal efficiency by about 12%, while decreasing permeate temperature slightly decreases the efficiency, whereas the feed salt concentration has a little effect on the thermal efficiency. Moreover, Al-Obaidani et al. (2008) indicated that the thermal efficiency increased with increasing feed temperature; feed flow rate, membrane thickness, and decreased as the solution concentration increases. A novel development reported by Li and Sirkar (2005) for VMD-based desalination found that the heat transfer efficiency was as high as ≥95%. Bandani et al. (1991) concluded that for DCMD the membrane properties (such as porosity, tortuosity and thermal conductivity) determine the thermal efficiency using pure water, where the effect of membrane

334 A.A. Al-Anezi et al.

thickness on the thermal efficiency is very peculiar. The thermal efficiency for pure water is independent of the membrane thickness and on the contrary, varies with the membrane thickness for salt solutions. Indeed, for more a thermally efficient MD system with acceptable permeate flux, there is always an optimum value for membrane thickness (Al-Obaidani et al., 2008; Bandini et al., 1991). According to theoretical investigations of AGMD, thermal efficiency in the range of 70% to 99% can be achieved by increasing the feed temperature to 90ºC and the temperature difference up to 55ºC respectively, where the air gap introduced to the MD module reduces the heat losses by conduction across the membrane to 1% to 30% (Banat and Simandl, 1998; Liu et al., 1998; Jönsson et al., 1985; Kubota et al., 1988). Guijt (Schneider et al., 1988) pointed out that the thermal efficiency of 85% to 90% was achieved experimentally by using hollow fibre module of air gap width 3 mm at feed temperature, and varies from 40ºC to 65ºC with small heat loss within the membrane module. Fane et al. (1987) reported that only 50% to 80% of the total heat transferred in the MD process is considered as latent heat, while the remaining 20 to 50% of the heat is lost by conduction. Schneider et al. (1988) and Jönsson et al. (1985) mentioned that using the membranes with higher porosity can reduce the thermal conductivity, which consequently, decreases the amount of heat lost by conduction, while, Martínez-Díez et al. (1999) concluded that both high feed temperature and high feed flow rate decreases the heat loss. Kimura et al. (1987) concluded that the heat conduction of the MD system is very small amount compared to the heat of vaporisation; therefore, it can be neglected. Different strategies have been suggested by Fane et al. (1987), Schofield et al. (1990), Andersson et al. (1985) and Khayet et al. (2006b) for minimising the heat lost by conduction across the membrane; the first, by using thicker membrane, i.e., increase membrane thickness, second, de-aerating the feed solution, which would enhance the latent heat, thirdly, working within the turbulent flow regime by using turbulence promoters, and finally, maintaining an air gap between the membrane and the condensation surface. It is worth stating that the MD process has the ability to recover latent heat of vaporisation for reuse, where the heat recovery minimises the heat requirement and enhances the operation cost (El-Bourawi et al., 2006; Bandini and Sarti, 2002). Schneider et al. (1988) mentioned that by using heat recovery, MD performance factors of up to 8 can be achieved. Heat recovery of up to 75% can be achieved by using a heat exchanger between the feed and warmed distillate, where the heat recovery depends on the heat exchanger area (Burgoyne and Vahdati, 2000). On the other hand, Kurokawa and Sawa (1996) pointed out that the increase in heat exchanger and membrane areas reduces the heat input. With regards to the membrane modules, larger effective membrane areas with lower flow rates both increase the contact time in the module, which gives rise to closer approach temperatures, and therefore more recoverable heat (Bandini and Sarti, 2002). Therefore, the heat exchanger capacity and membrane area should be optimised, in order to achieve the maximum heat recovery with an acceptable production (Ding et al., 2005).

7.2 Thermal energy consumption

In terms of energy consumption, Bui et al. (2010) used a simple energy balance for simulating and optimising DCMD to minimise the energy consumption for a wide range of operating conditions. They concluded that the highest thermal efficiency achieved was about 49.9%, while operating under the optimal operating conditions could reduce the total energy consumption up to 26.3%. Criscuoli et al. (2008) studied the heat

Potential of membrane distillation 335

requirement for different configurations of DCMD and VMD. They revealed that VMD performed better than DCMD, while the cross flow module is the most efficient design for achieving high flux and less energy consumption, where, the lowest energy consumption achieved was 1.1 (kWh/kg). Cabassud and Wirth (2003) performed an experimental investigation for VMD using hollow fibre modules for seawater desalination. They pointed out that the VMD could compete with RO in energy terms, but with a lower flux than that of RO. However, based on the computational analysis, VMD with energy consumption of less than 2 (kWh/m3) can be achieved using high permeable membranes with permeate flux range from 5 to 15 (L/m2 h). Moreover, if VMD is coupled with solar energy, and operated at high feed temperature, the flux could raise to 85 (L/m2 h) with energy consumption of 1.3 (kWh/m3), which could clearly compete with RO. On the other hand, an experimental study of AGMD for desalination was conducted by Gazagnes et al. (2007) using hydrophobic ceramic membranes, namely zirconia, alumina and alumino-silicate. They reported four times less energy consumption (1 kWh/m3) is required to obtain a permeate flux of 5 (L/m2 h) compared to RO. Additionally, the study consumes as much energy as a single pass or discontinuous VMD, which gives rise to flux one tenth lower compared to the study reported in Cabassud and Wirth (2003). The low energy consumption of a new AGMD technology, Memstill® with an average of 73.75 (MJ/m3) was reported, compared to 147.5 (MJ/m3) for MSF and MED. The Memstill® technology can use solar energy or low-grade heat with a temperature ranges from 50–100 to heat up the feed (Meindersma et al., 2006).

8 Economic in MD

From the economic standpoint, a preliminary economic investigation of MD process was carried out by Hanbury and Hodgkiess (1985), where, they noted that the MD process could become economically competitive for seawater desalination if the cost of membranes was reduced, and higher distillation heat transfer coefficients were achieved. In addition, Fane et al. (1987) carried out an economic analysis of producing 5,000 kg/h distilled water of a scaled up MD plant with heat recovery system. They deduced that the cost of produced water decreases with increasing feed temperature and the plant was shown to be cost competitive with RO. Al-Obaidani et al. (2008) conducted economic evaluation to assess the feasibility of MD as desalination process. They found that the estimated water cost from using the MD plant with heat recovery is $1.17 m3, where the cost of water is reduced to $0.64 m3 if MD is operated with low grade heat. The feasibility of solar thermal MD system for a capacity of 0.05 (m3/d) was carried out by Hogan et al. (1991). The system consisted of a hollow fibre module for MD and a heat recovery exchanger for reducing the capital cost, which was $3,500. They found that the heat exchangers are the most expensive items and with aid of a computer programme, they optimised heat recovery, solar collector area and membrane area to reduce the capital cost and achieve high flux. A comparative analysis in terms of energy and investment costs of the Memstill® technology between MD and RO for seawater desalination for a capacity of 105,000 (m3/day) was reported by Meindersma et al. (2006), where the heat supplied to the MD process was in cogeneration with electricity, or fuel fired, or from a waste heat source. The estimated total fixed cost for the MD process was $0.16–0.17/m3, while it was $0.25–0.35/m3 for RO. The MD process has

336 A.A. Al-Anezi et al.

been proved in its technical feasibility, but in terms of economic feasibility, further research is needed for optimising the parameters, where, utilising low-grade waste energy, integrating MD with other processes, and suitable heat recovery system would minimise the MD costs.

9 Fouling, scaling and flux decay

The fouling phenomenon is a major obstacle in conventional pressure driven membrane separation, while it is significantly less in MD process due to the relatively large membrane pores (Lawson and Lloyd, 1997). However, this phenomenon has not been thoroughly examined experimentally and theoretically in MD process (Lawson and Lloyd, 1997; l-Bourawi et al., 2006; Khayet and Mengual, 2004). The observed fouling, which may block the membrane pores, leads to flux decay, and subsequently a reduction in membrane life. The membrane surface characteristics, the hydrodynamic conditions and the chemical composition of feed liquid influence significantly the membrane fouling (Khayet and Mengual, 2004). The membrane fouling can be classified according to the fouling material as:

1 organic fouling caused by NOM

2 inorganic fouling due to the deposition on the membrane surface

3 biological fouling caused by microorganism growth on the membrane surface

4 scaling building up on the membrane surface due to the presence of high concentration solutions of minerals or salts (Lawson and Lloyd, 1997; Al-Amoudi, 2010).

Fouling is a common drawback of membrane processes, whereas the presence of deposition causes pore clogging in MD membranes and creates an additional layer on the membrane surface, which reduces membrane effective surface area for water vaporisation, and consequently, leads to flux decline. Moreover, increasing membrane fouling may cause pressure drop along membrane module, reducing flow rate, increasing the effect of temperature polarisation and consequently, reducing the permeate flux (Lawson and Lloyd, 1997; Kullab and Martin, 2011; Gryta, 2005). The scaling of saturation process solutions can cause both pore wetting and blocking, which is sometimes explained in the flux decay phenomenon (Lawson and Lloyd, 1997). The phenomenon of flux decay in MD was often observed in long-term operation (Curcio and Drioli, 2005) and may damage the MD module within a short period of time (Gryta, 2005). For feed concentration, membrane fouling can be a severe problem; therefore, an ultrafiltration pretreatment unit could be used for removing the larger particles that increase the viscosity of the stream through MD system (Calabro et al., 1994). Hsu et al. (2002) examined the effect of type of feed solutions on fouling of the MD process by using three different solutions:

1 raw seawater

2 raw water pretreated by microfiltration

3 3% NaCl.

Potential of membrane distillation 337

They pointed out that the permeate flux of the pretreated raw water increased by 25% and the permeate flux of feed solution of 3% NaCl was twice that of raw seawater. Moreover, Khayet et al. (Gryta, 2005) and Srisurichan et al. (2005) used pure water, then 0.1 M NaOH and/or HCl for cleaning a membrane of humic acid solutions containing NaCl, CaCl2. They found that the measured permeate flux was lower than that of the initial permeate flux by 6% and 14% respectively. Moreover, Mericq et al. (2010) observed a reduction of 8%, 11% and 24% on the initial permeate flux due to scaling of calcium crystals such as calcium carbonate CaCO3 and calcium sulphate CaSO4. Curcio et al. (2010) performed two steps cleaning of citric acid and sodium hydroxide for 20 min and they reported that the trans-membrane flux and membrane hydrophobicity had been completely restored. Hsu et al. (2002) used the ultrasonic irradiation technique for cleaning the membrane from fouling. Therefore, the feed pretreatment process and membrane cleaning are very important technique, for controlling membrane fouling and enhancing the permeate flux. In addition, membrane fouling and scaling in MD process can be effectively controlled by using appropriate hydrodynamic conditions, for example, lowering the feed temperature and increasing the feed flow rate can effectively reduce the speed of membrane fouling, while high saturation and supersaturating concentration salt solutions can drop the permeate flux sharply to zero over a short period of time. This is mainly attributed to the formation of salt crystals on the membrane surfaces (Yun et al., 2006; Tun et al., 2005; Gryta, 2010; Nghiem et al., 2011).

10 Long-term performance

The MD process was discovered in the late of 1960s. However, the investigations of long-term MD performance have received little attention. Schneider et al. (1998) conducted an experimental investigation of MD for producing pure water from normal tap water for a period of more than four months on a 24-hour basis. Moreover, fouling and scale formation were examined. They pointed out that decline in the permeate flux was observed after a few weeks of operation, whereupon the treatment using hydrochloric acid allowed the initial flux and the hydrophobicity of the membrane to be completely restored. However, during the last four weeks of the run, approximately 20% of the flux drop was recorded, through the acid treatment. Mericq et al. (2010) utilised VMD of seawater RO brines. For 6 to 8 h of experiments, they reported that no organic fouling or bio-fouling was observed, except scaling of calcium precipitation such as calcium carbonate and calcium sulphate which had very little effect on the permeate flux. In addition, Li and Sirkar (2004) observed a flux reduction of 23% from the initial flux during five days of continuous DCMD experiments for water desalination. Nghiem et al. (2011) mentioned that the higher flux decline during a continuous 1400 min of DCMD desalination of RO concentration can be attributed to the gradual adsorption of organic foulants into the surface of the PTFE membrane. In contrast, Izquierdo-Gil et al. (1999) concluded, from the experimental investigation of AGMD separation of sucrose aqueous solutions, that no significant changes in the permeate flux were observed during a period of a month, approximately, using the same membrane samples under the same experimental conditions. Two long operation AGMD experiments were performed for seawater desalination using two types of hydrophobic membranes (PVDF and PTFE). The first run was conducted uninterrupted for eight weeks, while the second was of

338 A.A. Al-Anezi et al.

six weeks duration. During the two months operation, the permeate flux for 1,500 h increased initially during 50 h, then declined for 160 h before reaching a steady state level. The scaling and fouling problems were not observed, except the membrane pore wetting because of an electrical shutdown at 1,075 h of the operation (Ding et al., 2006). DCMD experiments for the concentration and recovery of hydrochloric acid have investigated a capillary module of polypropylene (PP) membrane for more than one year. It has been mentioned that the hydrophobicity of the membrane was maintained, and the properties of the membrane did not change during the experiments (Tomaszewska et al., 1998). A new development for DCMD-based desalination experiments using a novel polypropylene hollow fibre membrane module was tested. A high stable water vapour flux (41–79 kg/m2 h) was achieved without any cleaning of the membrane module, and with the complete absence of the membrane pore wetting for over 400 h of continuous experiments (Banat and Simandl, 1994). The longest operational research on the DCMD module for producing pure water for over three years was conducted by Gryta (2005). The morphology of the hydrophobic capillary PP membranes remained unchanged using the SEM observations, and no membrane pore wetting was reported. However, the precipitation of CaCO3 was observed when tap water was used as a feed. Therefore, the capillary PP membranes used in DCMD investigations proved their thermal stability and the excellence characteristics of separation throughout the entire period of operation. In recent research, long-term, continuous DCMD experiments of up to 35 h were carried out on a semi-pilot plant for high salt concentrations at a feed temperature of 40ºC and a temperature gradient of 20ºC. The results showed that a significant drop in the system performance, and a flux decline of 45% of the original flux was observed after 35 h of operation (Curcio et al., 2010). Therefore, it can be concluded from the previous studies that more investigations are needed for long term MD performance, in order to obtain a quantitative measure of the flux decay, which was detected in some cases.

11 Conclusions

MD is a novel separation and purification process that promises numerous advantages. MD uses a hydrophobic or non-wetting, micro-porous membrane for separation. One side of the membrane is occupied by the liquid feed phase while on the other side the pure distillate condenses. The main separation parameters for the process are feed temperature, vapour pressure, and concentration of feed, specifically salt concentration in case of desalination. The unique feature of MD is that the driving force is temperature rather than pressure, as in other membrane separation techniques. Hydrophobicity of the membrane ensures the passage of water vapour only, restricting the entry of water completely. The feed water does not require any pretreatment in this method even for highly concentrated solutions and the high standards of distillate are obtained.

However the theoretical and experimental analysis of diffusion principle using mathematical models needs to be studied further. In depth studies of temperature and concentration polarisation on partial pressure of water vapour in membrane pores and on flux levels and the thermodynamic equilibrium relations of the unsaturated solution in the interface of membrane pores needs to be further studied. Furthermore, high requirements on the membrane material and membrane module must be satisfied in the applications of MD process to energy transformation. With the development of membrane materials and their manufacturing techniques, membrane materials and membrane modules with high

Potential of membrane distillation 339

performance cost ratio will be more specialised for various applications. In addition, MD will also be applied to more multi-field studies of different industries.

Energy and economy studies in MD are discussed in this article with respect to various configurations of MD. This article is a review paper for the MD process where the types of MD, membrane characteristics and its modules, experimental studies and the effect of operation parameters, energy economy, fouling, and long term performance were discussed.

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