Solar desalination by membrane distillation: Dispersion in energy ...

Desalination 308 (2013) 89–101

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Solar desalination by membrane distillation: Dispersion in energy consumption analysis and water production costs (a review) M. Khayet ⁎ Department of Applied Physics I, Faculty of Physics, University Complutense of Madrid, Avda. Complutense s/n 28040, Madrid, Spain

H I G H L I G H T S

G R A P H I C A L

► Report on economics, energy analysis and costs evaluations of MD technology. ► Dispersed and confusing water production costs (WPC) and energy consumption (EC) were reported. ► Useful equations and information on economics, energy consumption and water production costs are reported. ► More rigorous investigations and focused directions on economical analysis of MD systems should be conducted. ► A unified standard method for analysis and calculations should be followed to determine WPC of MD.

Cost elements needed to determine the WPC of MD process.

a r t i c l e

a b s t r a c t

i n f o

Article history: Received 20 March 2012 Received in revised form 4 July 2012 Accepted 13 July 2012 Available online 5 August 2012 Keywords: Membrane distillation Desalination Gained output ratio Energy consumption Water production cost Solar energy

A B S T R A C T

The non-isothermal membrane distillation (MD) separation process is known for about 50 years and very few studies are reported on its economics, energy analysis and costs evaluations. Dispersed and confusing water production costs (WPC) and specific energy consumption (EC) analysis were reported. Most of them are simulated and others are based on various costs assumptions. At present, the common asked questions about the published papers in MD including EC and WPC are: how these reported calculations on WPC and EC were made?, what is the current WPC of MD?, and how WPC of MD can be improved?. An overview of most studies carried out on these issues is presented and some useful equations and information in this context are reported. Comparison to other separation processes used in desalination is made. At present, the main challenge for large-scale MD is EC and WPC. New directions on MD should be raised. More rigorous investigations and focused directions on economical analysis of MD systems should be conducted. A unified standard method for analysis and calculations should be followed to determine WPC. For the benefit of MD process, one should be cautious when reporting simulated, non-realistic and non-contrasted WPC. © 2012 Elsevier B.V. All rights reserved.

1. Introduction Over the years various solar desalination technologies have been developed including thermal distillation (i.e. multistage flash, MSF;

⁎ Tel.: +34 91 3945185; fax: +34 91 3945191. E-mail address: khayetm@fis.ucm.es. 0011-9164/$ – see front matter © 2012 Elsevier B.V. All rights reserved. doi:10.1016/j.desal.2012.07.010

multiple effect evaporation, ME; vapor compression, VC) and membrane processes (i.e. reverse osmosis, RO; electrodialysis, ED; nanofiltration, NF). Currently, the major technologies dominating the market for both seawater and brackish water desalination are MSF and RO. This last process accounts for over 50% of the installed capacity [1,2]. It is to point out that desalination by any type of technology is commonly considered a capital and energy intensive process [1,2]. In fact, the theoretical lower bound of the specific energy needed for

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M. Khayet / Desalination 308 (2013) 89–101

desalination was estimated from the free energy-change and found to be 0.7 kWh/m3 of fresh water produced without energy recovery, while for a given energy recovery higher than zero per cent this theoretical minimum value is higher (i.e. 0.81 kWh/m 3 for 25% energy recovery, 0.97 kWh/m 3 for 50% energy recovery and 1.29 kWh/m 3 for 75% energy recovery) [2,3]. Moreover, as the salt concentration increases, the minimum energy required for desalination also increases. However, the actual specific energy consumption (EC) is larger than these cited values. For example, taking into consideration all recent advances on energy efficiency of RO desalination, the specific energy of desalination has been reduced to a value near 1.8 kWh/m 3, which possibly may decrease further with more improvements in membrane performance and very efficient energy recovery systems [4]. In this context, renewable energy is being used as one of the possible solutions to tackle the energy problem [5]. Energy and cost analysis in thermal distillation (ME) and membrane desalination processes MSF and RO are well studied; however only a little information is available on energy analysis and cost estimations for membrane distillation (MD) process [6,7]. Few studies were reported on the economics, energy analysis and costs evaluations [8–37]. One of the reasons is that MD is not yet fully applied for commercial scale and the capital investment costs that include membranes and modules are fluctuating. Other cost-related information such as pretreatments, optimum flow conditions, long-term MD performance, fouling, and membrane life are not yet available at a satisfying level. The highly concentrated brine discharged by the process is also of ecological concern and its related cost should be taken into consideration. In some reported MD papers [8,10,13,15–21,23,25–27,32–34], authors dedicated only a small section to the energy consumption and/or water production cost (WPC) without specifying the followed calculations of the costs and energy analysis. There is no full agreement on standard for calculations of WPC. Moreover, although MD process is used in different applications fields not only in desalination, up to date practically all authors reported the energy consumption and cost evaluation of MD systems used only for the treatment of seawater and brackish water. Different and even conflicting conclusions are drawn depending on the laboratory system or the pilot plant from which the data for the analysis was produced. For example, this paper will show that the reported EC and WPC in MD vary widely, depending on the authors and the configuration used, from about 1 to 9,000 kWh/m 3 for EC (3 orders of magnitude dispersion) or even higher and from $0.3/m 3 to $130/m 3 (4 orders of magnitude fluctuations) for WPC. Referenced details will be presented. Taking into account that MD exhibits the advantage of using renewable energy sources such as solar and geothermal energy or any low-grade and industrial waste heat, practically in all reports there is one common conclusion, i.e. the WPC can go down considerably if inexpensive heat source is available. As it will be shown in this paper, it was reported that WPC would be as low as $0.64/m 3 if waste heat is used. According to some other reports, if solar energy systems were considered the capital cost (CC) would increase considerably due to the enhancement of the capital investment costs (solar systems and monitoring equipments). In fact, for an optimum MD plant design, not only renewable energy sources are interesting from point of view of reduction of long‐term WPC but also improved energy recovery devices and hybrid systems should be employed. In this paper, comparisons with other processes such as RO and MSF are also made, although these processes are well established industrially under strict optimization, whereas MD process is not. Optimized MD plants must be designed and developed first and then rigorous cost and economic analysis for comparison will become possible. More intensive and focused MD research efforts in this field are needed, not only theoretically but also experimentally looking forward the decrease of EC and WPC.

2. Dispersions in energy efficiency and specific energy consumption Unlike the pressure-driven membrane separation processes such as RO and NF, energy consumption in MD systems includes the thermal energy necessary to heat up the feed aqueous solution to be treated and to cool down the permeate aqueous solution or condensation in heat exchangers and the electrical energy required to run the circulation pumps, vacuum pumps or compressors depending on the MD configuration. The effects of different MD operating parameters and membrane module designs on the EC are barely studied. In fact, EC of pilot MD systems has been performed mainly for fixed operating conditions, although these are not optimized [13,14,17,18,21,22,26,27,29,30]. In most studies, authors reported only on the thermal efficiency (εT) of their MD systems or membrane modules [6,7,30,38–41]. The εT was defined as the ratio between the amount of heat brought into the membrane and the heat actually used for evaporation of the feed to produce fresh water: εT ¼

J w AΔH v;w QV ¼ Qm QV þ QC

ð1Þ

where Jw is the permeate flux, A is the membrane area, ΔHv,w is the enthalpy of evaporation of water, QC is the heat transferred by conduction through the membrane material and the gas-filled pores of the membrane (i.e. heat lost in MD), QV is the latent heat associated with the vaporized molecules that lead to the permeate flux, and Qm is the total heat flux transferred through the membrane. In most cases, at steady state conditions, Qm is considered equal to the heat transfer in the feed channel of the membrane module (Qf) using the following equation [42]: _ f cp T f;in −T f;out Qf ¼ m

ð2Þ

_ f is the feed flow rate, cp is the specific heat of the feed soluwhere m tion, and Tf,in and Tf,out are the temperatures of the feed solution at the inlet and outlet of the membrane module, respectively. The effects of different MD operating conditions on εT were studied [30,39–41]. It was observed that εT is very sensitive to the feed temperature [30,39]. This is mainly due to the exponentially increased mass flux (Jw) and to the decrease of the amount of heat lost by conduction (QC) through the membrane with the increase of the feed temperature. At high feed temperatures the heat transferred through the membrane by conduction will be negligible compared to the heat transferred due to the mass flux. Therefore, the energy consumption per unit of distillate may be reduced appreciably at high operating feed temperatures. In order to increase εT in MD, various possibilities were adopted both in membranes and modules designs (i.e. optimum membrane thickness, high porosity, adequate membrane materials, etc.) [6,7]. Very few authors do believe that it is more adequate to use energy efficiency (εE) to characterize an MD system instead of the thermal efficiency (εT), since energy efficiency takes into consideration the global energy input (Ein), which includes both thermal energy (Et) and electrical energy (Ee) [7,35]: εE ¼

J w AΔH v;w J w AΔH v;w ¼ Ein Et þ Ee

ð3Þ

Different attempts were made looking for the reduction of εE including the use of different heat recovery systems and membrane modules with different designs, heat turbulence promoters to decrease the temperature polarization effect, adequate insulation materials, pipes and other plant accessories in order to limit heat loss to environment, etc. [9,10,16,18–20,25–27,30,33]. Recently, to estimate the energy efficiency of MD process, the concept of gained output ratio (GOR), which is defined also as the ratio of heat associated with mass transfer to the energy input (Eq. (3)), has


been applied. The GOR reflects how well the energy input in the system is utilized for the water production. The higher the GOR value is the better is the performance of the system. In MD process the GOR is also a dimensionless parameter defined as the ratio of the latent heat of evaporation of the produced water to the total input energy in the MD system: GOR ¼ εE ¼

J w AΔH v;w Ein

ð4Þ

The following equation that considers the heat recovery factor (HR) in the MD system was also considered for the GOR calculation [24]: GOR ¼ εT H R ¼ εT

ΔT MD;module ΔT MD þ ΔT HE

ð5Þ

where HR refers to the number of energy uses and can be defined differently depending on the arrangements of the MD system. Fane et al. [9] defined it as: HR ¼

ΔT MD;module ΔT MD þ ΔT HE

ð6Þ

where ΔTMD,module is the axial temperature drop along the MD module, ΔTMD is the temperature difference across the membrane and ΔTHE is the temperature difference across the heat exchanger(s). Hogan et al. [10] defined HR as the maximum heat recoverable in the main heat recovery exchanger (QHE) divided by the heat transferred in the membrane module (QMD): HR ¼

Q HE Q MD

ð7Þ

High HR values were achieved when using high membrane areas and low flow rates because of the large residence time of the liquids in the membrane module. However, the permeate flux (i.e. water production per unit of membrane area) decreased when increasing the membrane area and when applying low flow rates because of the temperature and concentration polarization effects. Therefore, the design and use of heat recovery systems should be optimized [9,30,43]. Moreover, since in a single-stage MD system the rejected brine (i.e. retentate) exits at a temperature higher than the feed inlet temperature, to improve energy efficiency in MD, this energy could be recovered to preheat the feed aqueous solution using efficient heat exchangers so that the process heat requirement is reduced. This is also possible through multistage MD systems [24,44]. The maximum heat recovery should be as high as possible. However, the installment of heat recovery systems increases the capital investment costs. Gilron et al. [24] reported a tradeoff between the theoretical

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GOR, the DCMD permeate flux and the number of stages suggesting the presence of an optimum number. The theoretical GOR increased from about 2.5 to 10, whereas the permeate flux decreased from 12 to 2 kg/m2.h with the increase of the number of stages in cascades from 2 to 11. It is worth quoting that few research studies have been focused on the effects of the membrane module design and MD operating parameters on the GOR [9,24,36,45]. In fact, GOR value is less than unity for simple effect systems and higher than unity for multi-effect systems. For example, a solar still system commonly has a GOR value less than 1, whereas good multi-effect distillation (MED) systems exhibit GOR values as high as 12 [46]. Recently, Lee et al. [45] by means of a numerical simulation predicted also a GOR value of 12 for a countercurrent cascade of cross-flow DCMD modules in 10 stages. Table 1 shows as an example the dispersion of the GOR values of some tested MD systems in the range 0.3–8.1. As can be seen in Table 1 the GOR does not depend on the MD configuration and most experimental MD systems have GOR values less than unity indicating low MD performance (i.e. poor heat recovery). This is mainly because these MD systems are operated in a single stage. However, the combined one stage MED/DCMD hybrid system working at atmospheric pressure achieved a GOR value of 4.2 [47]. In this hybrid system the relatively hot brine and distilled water of the MED of 1 m 3 volume are reused as feed and permeate in shell-and-tube DCMD modules. The MED alone exhibited a GOR of 3.7 and a water production of 16 kg/h at a temperature of 85 °C and a circulation feed solution of 170 kg/h. The MED/MD hybrid system permitted an increase of the production by about 7.5%, with an electrical conductivity of about 12 μS/cm and an improvement of the GOR by practically 10%. Solar-powered liquid and air gap MD desalination plants with capacities up to 10 m3/day was developed by Fraunhofer ISE (Germany) using a spiral wound membrane module with heat recovery [16,25–27,33,37,48]. Details of the different pilot plants may be found in [16,25–27,33,37,48]. In these pilot plants the feed aqueous solution to be treated passes first through the condenser channel and is gradually warmed by the latent heat of condensation. When it comes out of the condenser it is fed to the solar thermal flat collector for further heating before contacting the membrane. Among the reported characteristics of the membrane modules were the effective membrane area of 7–12 m2, and the GOR values of about 3–6 [33]. For example, Koschikowski et al. [16], by using a module with an effective membrane area of 8 m2 in a compact plant, reported a GOR value of 5.5 at 350 l/h flow rate and 75 °C evaporator inlet temperature. However, when using a similar compact pilot plant with 10 m2 membrane area operated with brackish water, AGMD permeate fluxes as high as 120 l/day with an electrical conductivity of about 5 μS/cm were obtained and the GOR values were found to be lower than unity (i.e. 0.3–0.9) [25]. Furthermore, when using a larger pilot plant with heat storage tank and 4 membrane modules with a total area of 40 m2 operated with untreated (i.e. without chemical treatment

Table 1 GOR values of tested MD systems. GOR

Year

Observations

Ref.

4.2 5.5 1.2–8.1 Average: 2.6±1.1 0.3–0.9 0.4–0.7 0.8–5.6 Average: 2.1±0.6 3–6

1998 2003 2005

Integrated MED/DCMD system, water production: 12.4 kg/h. Compact AGMD solar-driven plant, membrane area: 8 m2, distillate: ≈20 l/h, EC: 117 kWh/m3. Liquid gap MD spiral wound module used in compact solar-driven plant, membrane area: 8.5–10 m2, distillate: 47 ± 20 l/day, recovery ratio: 44 ± 5, EC: 68.8–492 kWh/m3. Compact AGMD solar-driven, membrane area: 10 m2, water production up to 120 l/day, EC: 200–300 kwh/m3. Two loop AGMD solar-driven plant, membrane area: 40 m2, permeate flux: 1.5 l/m2.h, seawater from Red Sea as feed, EC: 200–300 kwh/m3. Liquid gap MD spiral wound module used in compact solar-driven plant, membrane area: 8.5–10 m2, distillate: 44 ± 190 l/day, recovery ratio: 42 ± 0.9, EC: 203.3–441.4 kWh/m3. AGMD spiral wound module used in compact and large solar-driven plants, membrane area: 7–12 m2, distillate: 10–30 l/h, EC: 100–200 kWh/m3. Liquid gap MD spiral wound module used in compact solar-driven plant, membrane area: 8.5–10 m2, distillate: 76 ± 22 l/day, recovery ratio: 18 ± 14, EC: 138.5–499.1 kWh/m3. Cascade of cross-flow DCMD, membrane area: 0.4 m2 (per stage), feed flow rate: 0.5 l/min (0.008 kg/s), 8 stages in series.

[47] [16] [48]

0.5–5.1 Average: 3.4±0.5 4.1

2007 2007 2008 2009 2009/2010 2011

[25] [26] [48] [33] [48] [45]

92


usually used in RO) seawater from red sea (Jordan), fresh water production varying from 5 to 27 l/m2.h with electrical conductivity values varying between 20 and 250 μS/cm (i.e. average salt rejection factor of 98%) were produced [26]. In this case, the calculated GOR values were also found to be lower than unity, in the range 0.4–0.7. It may be expected from Eqs.1, 3, and 4 that larger membrane areas lead to higher GOR values because of the increase of the water production. However, as reported previously this is not the case for the solar desalination systems developed by Fraunhofer ISE (Germany) [16,25–27,33]. Theoretical GOR analysis was also carried out based on experimentally tests MD systems [36,49]. Recently, Zuo et al. [36] simulated a DCMD system using Aspen Plus with a special focus on energy efficiency and economic analyses. A cross-flow membrane module previously presented by Song et al. [50] was considered. The relationships between the GOR and the membrane area as well as the DCMD operating parameters were investigated. It was observed that a critical membrane area did exist (i.e. 4 m 2) below which significant enhancements of the GOR values were detected with increasing the membrane area. Over this critical membrane area, no significant effect of the membrane area on the GOR was observed. In fact, more capital costs are required when using larger membrane modules. By increasing the feed temperature and the feed and permeate flow rates the GOR increases. However, above certain levels of feed and permeate flow rates the GOR became insensitive to the flow rates because the electrical energy input for pumping became important. It was concluded that energy recovery was very important for the operation of DCMD system. Nevertheless, the highest achieved GOR theoretical value was 1.4 reported for a feed temperature of 90 °C, a membrane area of 10 m 2, a feed velocity of 0.04 m/s, a permeate velocity of 0.48 m/s, a permeate temperature of 25 °C and a feed concentration of 3 wt.% sodium chloride (model seawater).

Lee et al. [45] performed both numerical simulations and experimental studies integrating a countercurrent cascade of cross-flow DCMD modules in conjunction with and without heat recovery devices. The effects of different operating parameters such as the number of stages, permeate and brine inlet temperatures, and inlet flow rates of the brine and distillate were studied. Over the range of the experimental conditions studied, the thermal efficiency of the membrane modules was found to be between 0.73 and 0.87 and the GOR reached a value of about 4.1 for 8 DCMD modules in a cascade, a brine temperature of 90 °C and equal inlet flow rates of the brine and the distillate of 0.5 l/min. The GOR increased with the number of stages in a cascade and without heat recovery it is lower than 1 independently of the number of modules. It was reported an increase of the simulated GOR value with the increase of the distillate temperature maintaining the brine temperature fixed and with the increase of the brine temperature maintaining the distillate temperature fixed. Summers et al. [49] by performing analytical models for the three configurations VMD, AGMD and DCMD reported that the membrane module geometry and its design exerted an important effect on the predicted GOR, particularly its effective length in the case of AGMD and DCMD configurations. For example, for a membrane module length of 140 m the achieved theoretical GOR value was 5 for DCMD and about 5.7 for AGMD. However, the theoretical GOR of VMD systems was always lower than 1 even when heat recovery by recirculation of the rejected brine was considered. For AGMD, the most significant parameter affecting the theoretical GOR was the air gap width. Another interesting characteristic parameter for a desalination plant is the specific energy consumption (EC) defined as the energy input required to produce 1 m 3 of distillate (i.e. ratio of energy supplied to the volume of produced fresh water). Table 2 summarizes

Table 2 Estimated specific energy consumption, EC, of different MD systems. EC (kWh/m3)

Year

Observations

Ref.

1.25 13 15 2.25 2.58 1.2 3.2 117 140–200 30.8 68.8–492

1983 1999 1999 1999 1999 2003 2003 2003 2003 2005 2005

[8] [14] [14] [14] [14] [15] [15] [16] [16] [17] [48]

20.5–66.7 ≈1 2.05 ≈1.63 28.0 ≈27.54 200–300 200–300 203.3–441.4

2006 2007 2007 2007 2007 2007 2007 2007 2008

138.5–499.1

2009/ 2010

AGMD, use of waste energy, feed temperature b70 °C, water production: 5 m3/day. NF/RO/DCMD hybrid system. RO/DCMD hybrid system. RO/DCMD hybrid system, thermal energy available in the plant or stream. NF/RO/DCMD hybrid system, thermal energy available in the plant or stream. VMD discontinuous flow, permeate flux: 0.5–0.7 l/m2.h. VMD single-pass flow, permeate flux: 0.7 l/m2.h. Compact AGMD solar-driven plant, membrane area: 8 m2, distillate: ≈20 l/h (feed temperature of 75 °C). Compact AGMD solar-driven plant, membrane area: 8 m2, distillate: 20–30 l/h, feed temperature: 60–85 °C, GOR: 4–6. AGMD, use of sensible heat of geothermal Liquid gap MD spiral wound module used in compact solar-driven plant, membrane area: 8.5–10 m2, distillate: 47±20 l/day, recovery ratio: 44±5, GOR: 1.2–8.1 (average: 2.6±1.1). Memstill® units, water production: 25–50 m3/day per module. AGMD, permeate flux: 5.2 l/m2.h. MD hybrid system (MF/NF/MCr/RO/MD) without energy recovery device, thermal energy available in the plant or stream. MD hybrid system (MF/NF/MCr/RO/MD) with energy recovery device, thermal energy available in the plant or stream. MD hybrid system (MF/NF/MCr/RO/MD) without energy recovery device. MD hybrid system (MF/NF/MCr/RO/MD) with energy recovery device. Compact AGMD solar-driven plant, membrane area: 10 m2, water production up to 120 l/day. Two loop AGMD solar-driven plant, membrane area: 40 m2, permeate flux: 144–792 l/day, seawater from Red Sea as feed. Liquid gap MD spiral wound module used in compact solar-driven plant, membrane area: 8.5–10 m2, distillate: 44±190 l/day, recovery ratio: 42±0.9, GOR: 0.8–5.6 (average: 2.1±0.6) Liquid gap MD spiral wound module used in compact solar-driven plant, membrane area: 8.5–10 m2, distillate: 76±22 l/day, recovery ratio: 18±14, GOR: 0.5–5.1 (average: 3.4±0.5).

EC higher than 500 kWh/m3 600–1600 1999 870–1550 2006 3550–4580

2008

1080–2980

2008

8100.8–9079.5

2009

DCMD spiral wound module, membrane area: 4 m2, permeate flux: 3–8 l/m2.h, feed temperature: 47–80 °C. DCMD spiral wound module, membrane area: 4 m2, water production: 12–39.86 l/h, feed temperature: 47.7–75 °C, distillate temperature: 17.0–26.8 °C. DCMD (Longitudinal flow membrane module), permeate flux: 7.81–25.4 kg/m2.h, feed temperature: 39.8–59 °C, distillate temperature: 13.4–14.3 °C. VMD (cross-flow membrane module), permeate flux: 29.7–56.2 kg/m2.h, feed temperature: 53.9–59.2 °C, vacuum pressure: 10–60 mbar vacuum pressure. VMD, membrane area: 0.09 m2, permeate flux: 4.43–29.75 kg/m2.h, thermal and electrical energy consumption (circulation power and vacuum pump, heat): 4.43–29.75 kW.

[20] [21] [22] [22] [22] [22] [25] [26] [48] [48]

[13] [18] [31] [31] [34]


the estimated EC reported for different MD systems and pilot plants. As can be seen EC in MD fluctuates in 3 orders of magnitude from about 1 up to 9,000 kWh/m 3. Carlsson [8] stated that the EC of a MD system can be as low as 1.25 kWh/m3, the standard modules in large‐scale MD systems would produce about 5 m 3/day per module and for large‐scale plants the cost would be reduced. However, no analytical details of the MD system and the operating parameters were reported. Koschikowski et al. [16], by using an AGMD spiral wound module with an effective membrane area of 8 m 2 in a compact solar-driven desalination plant, reported an EC of about 117 kWh/m 3 with a GOR value of 5.5 at 350 l/h flow rate and 75 °C evaporator inlet temperature. EC ranging from 140 and 200 kWh/m 3 with GOR values between 4 and 6 were also claimed in the same paper [16]. With a similar compact pilot plant having higher membrane area (10 m 2) operated with brackish water, the AGMD permeate flux was as high as 120 l/day with an electrical conductivity of about 5 μS/cm, and the thermal energy required was in the range 200 and 300 kWh/m 3 [25]. For a larger pilot plant with a total area of 40 m 2 operated with untreated seawater from red sea (Jordan), the water production was varied from 5 to 27 l/m 2.h with an electrical conductivity varying between 20 and 250 μS/cm and the estimated EC was found to be between 200 and 300 kWh/m 3 with GOR values between 0.4 and 0.7 [26]. Recently, Winter et al. [37] studied the effects of the membrane area of spiral wound liquid gap MD modules (LGMD, Fraunhofer ISE) on the EC. The studied membrane areas were 5, 10 and 14 m 2 and the considered energy consumption was only the thermal. The channel height of the module was maintained fixed, whereas the channel length of the module was varied. The hot feed and the condensation temperatures at the module inlets were maintained fixed at 80 and 25 °C, respectively; the feed flow rate was changed from 200 to 500 kg/h and distilled water was used as feed. It was observed that the increase of the membrane surface area did not affect too much the total permeate flux, but it significantly reduced the EC. For example, for a feed flow rate of 400 kg/h an increase of the membrane area from 5 to 14 m 2 resulted in an increase of the permeate flux of only 8.5% and a decrease of the EC of 57.1%. However, when a salt solution (50 g/L) was used as feed, it was observed a decrease of both the permeate flux and the EC with the increase of the module membrane area. The reduction of the EC was more pronounced for higher feed flow rates. It was concluded that the membrane module design depended not only on the MD operation conditions, but also on the concentration of the feed solution and the module cost, and the EC was in the range 130–207 kWh/m 3. For a similar solar-driven desalination plant composed of a spiral wound LGMD module (Fraunhofer ISE) having a membrane area of 8.5–10 m 2, EC values between 68.8 and 499.1 kWh/m 3 were reported with water production in the range 1–117 l/day, recovery ratio in the range 0.1–55% and GOR values between 0.5 and 8.1 [48]. Memstill® technology based also on AGMD configuration with heat recovery was proposed by the research and technology organization TNO (Netherlands Organization for Applied Scientific Research) [19,20]. A cold saline water flows through a condenser with non-permeable walls, increasing its temperature progressively due to the condensing permeate, and then passes through a heat exchanger where additional heat is added before entering in direct contact with the membrane in counter current flow configuration. The average energy consumption was claimed to be 73.75 MJ/m3. When the considered heat supply to the Memstill system was fuel fired or generated by cogeneration of heat and electricity, the EC was found to be 64.9 kWh/m3 (i.e. 231 MJ/m3 heat consumption plus 0.75 kWh/m3 electric energy supplied), whereas in the case of waste heat source the EC was lower, 39.4 kWh/m 3 (i.e. 139 MJ/m 3 heat consumption plus 0.75 kWh/m 3 electric energy supplied) [20]. By using sensible heat of geothermal water and AGMD system, Bouguecha et al. [17] obtained an EC of 30.83 kWh/m 3 (i.e. 111 kJ/kg).

93

In this case, a pretreatment by a conical fluidized bed crystallizer was used because of precipitation of calcium carbonate (CaCO3) that took place due to the decrease of pressure and loss of carbon dioxide (CO2) when water in the well moved close the surface. Three AGMD modules associated in series were considered. Each module consisted of two cells arranged in parallel configuration. The electrical conductivity of the permeate was constant at 6 μS/cm and the overall water conversion reached 25% (i.e. 10% in the first stage, 9% in the second stage and 6% in the third stage). The theoretical value of the EC was found to be 11.1–13.9 kWh/m3 (i.e. 40–50 kJ/kg), which is lower than the experimental one. The difference between the experimental and theoretical values of the EC was attributed by the authors to the exponential variation of the permeate flux with the temperature and also to the fact that the second and third stages of AGMD coupled to geothermal resources were operated at low temperatures leading to lower permeate fluxes and lower efficiencies. However, details of the calculations, membrane type, some of the operating AGMD conditions as well as geothermal water characteristics and real coupling were not reported [17]. Gazagnes et al. [21] indicated that AGMD process is able to produce a permeate flux of 5.2 l/m 2.h with an EC of about 1 kWh/m 3. This value is quite similar to the theoretical lower bound of the specific energy needed for desalination (see Introduction section). However, detailed discussions with more rigorous calculations were not reported in order to believe in the reported EC number. In fact, higher EC values than those reported in Table 2 were also estimated in some MD studies [13,18,31,34]. Khayet et al. [18] studied DCMD in a pilot plant designed with heat recovery systems and a spiral wound polyetetrafluoroethylene (PTFE) membrane module of an effective membrane area of 4 m 2. The pilot plant has three heat exchangers for heat recovery (i.e. the distillate transfers the heat to the retentate stream). The DCMD experiments showed that the best conditions for running the process were at higher inlet feed temperatures (70–80 °C). It was observed that the EC decreased with the increase of the inlet feed temperature. For example, at a feed temperature ranging from 80 to 85 °C, a feed flow rate of 1380 l/h, a permeate flow rate of 1260 l/h and a permeate temperature varying from 16.5 to 26.8 °C, the EC was about 600 kWh/m 3. This value is higher than those reported in the first group of Table 2. For high feed temperatures, the permeate flux is high and therefore the EC is low partly due to the heat recovery. Previously, Zakrzewska-Trznadel et al. [13] carried out DCMD tests using the same pilot plant and observed that the increased of the flow rate of cooling water caused a slight increase of the permeate flux but it induced an enhancement of the EC. Wang et al. [34] proposed a small solar VMD hybrid experimental system for potable water production from underground water. A polypropylene (PP) hollow fiber membrane module with a total membrane area of 0.09 m 2, an external condenser and a vacuum pump were employed. It was observed that the power consumption of the vacuum pump (≈0.18 kW) and the feed circulation pump (≈0.37 kW) are much lower than the heat power consumption, especially for high permeate fluxes. A high water production of 29.75 kg/m 2.h with a total power consumption of 21.69 kW was reported. This means that the EC is 8100.8 kWh/m3. A higher EC value can be calculated for a lower water production rate of 3.07 kg/m2.h with a total power consumption of 3.62 kW. Surprisingly, compared to the values of the EC reported in the first groups of Table 2, the obtained values by Wang et al. [34] are more than two orders of magnitude greater. One of the reasons of the observed high EC values is the small membrane module used without considering heat recovery systems. High EC for laboratory MD systems may be expected. Evaluation of energy requirements in DCMD and VMD laboratories systems have been made by Criscuoli et al. [31] using plate-and-frame membrane modules of about 40 cm2 effective membrane area. Longitudinal and transversal flow configurations were studied and different MD operating conditions were tested. The lowest obtained values of the EC were 3546.3 kWh/m 3 for longitudinal flow in DCMD and

94


1108.4 kWh/m 3 for cross-flow VMD membrane module. In terms of permeate flux, EC and thermal efficiency VMD performance was found to be better than DCMD configuration. In some studies [15,28,35] authors developed mathematical models trying to evaluate the EC of MD and to study the effects of different operating conditions on EC. Based on a developed computational analysis using a theoretical model describing the coupled heat and mass transfer, Cabassud and Wirth [15] studied the EC in two VMD systems, a continuous single-pass VMD system and a discontinuous VMD operation in which the retentate was returned back to the feed container. However, theoretical model specifications were not reported. For the continuous single-pass system the EC was estimated to be 3.2 kWh/m 3, while for the discontinuous system the EC decreased to 1.2 kWh/m 3, which was practically the energy required by the vacuum pump. For VMD process, when the feed temperature is as low as the ambient temperature only the energy consumed by the vacuum pump and the feed circulation pump are necessary to be considered. At 25 °C, Cabassud and Wirth [15] found that the VMD permeate flux was independent on the Reynolds number (Re) and the minimum value of the EC was 1.2 kWh/m 3 at a Re number of 2000 when the feed salt concentration was 30 g/l and the downstream pressure was 100 Pa. Under these conditions, the VMD permeate flux was 13 l/m 2.h. On the other hand, when the feed concentration was 300 g/l, the minimum EC was 1.3 kWh/m 3 at the Re number of 500 with a VMD permeate flux of 8.8 l/m 2 h. For higher feed temperatures than the ambient temperature, the main energy consumption will be for heating the aqueous feed solution. This heat energy was found to be as high as 100 kWh/m 3; however, the vacuum energy consumption was as low as 1.3 kWh/m 3 and almost negligible as compared to the heating energy consumption [15]. In contrast, the energy consumption for feed circulation increased with the increase of the flow rate. Cabassud and Wirth [15] indicated that the energy required for feed circulation was negligible compared to the vacuum pump energy until a salt concentration of 300 g/l and a Re number of 7000 are reached. The authors further stated that for a membrane area 100 times larger than that of the experimental membranes and at low temperatures of about 25 °C, VMD can compete with RO on EC (b2 kWh/m3) with approximately the same water production. Alklaibi [28] developed a mathematical model for a stand-alone MD process with heat recovery trying to relate the energy required for heating and pumping to the heat and mass transfer in a MD system. Most of the heat required to increase the temperature of the feed aqueous solution to the inlet temperature of the hot channel of the MD module is recovered by a heat exchanger. The calculated values of the EC for different membrane permeabilities and feed temperatures were found to be between 6 and 1500 kWh/m 3. The obtained EC was reduced significantly by using a membrane with a high permeability and by discharging the feed at the outlet of the MD module at a high temperature. For example, for a feed temperature at the outlet of the membrane module of 25 °C, when the membrane permeability was increased 10 times, the calculated EC was reduced more than 3 times. However, for a feed temperature at the outlet of the membrane module of 60 °C, the EC was decreased more than 5 times for the same increase of the membrane permeability. Bui et al. [35] developed also a simulation procedure to optimize DCMD and to minimize its EC. This energy accounted for the heat exchanged between the feed and permeate liquids within the membrane module and the power for their pumping. It was found that operating the DCMD process under the optimal conditions could result in a reduction in the EC of up to 26.3%. Another parameter used in MD to evaluate EC output is “exergy” [14,22,30,51]. In thermodynamics, the exergy of a system is the maximum useful work possible during a process that brings the system into equilibrium with the surroundings. It represents the useful part of energy for a system in its environment. After the system and surroundings reach equilibrium, the exergy is zero. In other words, the

total energy is divided into two parts, one anergy and the other exergy (or available work). The anergy is the part of energy that is forced to be given to the environment as heat in conditions of complete degradation. For an isothermal process, exergy and energy are interchangeable terms, and there is no anergy. The exergy is the part of energy that is convertible into all other forms of energy. For a system in which the governing parameters are the temperature, pressure and composition, the exergy of a flow stream, EX, was written as [30]. T

P

C

EX ¼ EX þ EX þ EX

ð8Þ

where EXT, EXP and EXC are the temperature, pressure and concentration exergy terms, respectively, that were defined as follows: T T _ cp ðT−T 0 Þ−T 0 ln EX ¼ m T0 P _ EX ¼ m

P−P 0 ρ

C _ cp ðns RT 0 lnðxs ÞÞ EX ¼ −m

ð9Þ

ð10Þ ð11Þ

_ is the mass flow rate, cp is its specific heat, ρ is the density, ns where m is the solvent concentration (mol/kg solution), xs is the solvent mass fraction and the subscript 0 stands for reference state. The indicator of how much exergy remains undestroyed during the operation of the system was termed exergetic efficiency and was defined as [22,30]: ξ¼

Exout 100 Exin

ð12Þ

where Exout and Exin indicate exergy input and exergy output to the system, respectively. Criscuoli and Drioli [14] performed energetic and exergetic analysis of MD hybrid systems, RO/MD system (i.e. RO brine was treated by MD) and NF/RO/MD system (i.e. NF used as pretreatment step for the RO/MD system). The obtained EC was 15 kWh/m 3 for the RO/MD hybrid system, whereas that of the NF/RO/MD hybrid system was found to be lower, 13 kWh/m 3. In the case of available thermal energy in the plants, smaller EC values, 2.25 kWh/m 3 and 2.58 kWh/m3, were reported for the RO/MD hybrid system and the NF/RO/MD hybrid system, respectively. Macedonio et al. [22] also made energetic and exergetic analysis as well as economic evaluation for seawater desalination by integrated membrane systems. One of the systems combined microfiltration (MF), nanofiltration (NF), membrane crystallization (MCr), RO and MD in one hybrid system. MCr was applied on NF retentate and MD was applied on RO retentate. Without energy recovery the EC was found to be 28 kWh/m 3. When the throttling valves on the brine stream were replaced by an energy recovery system such as a Pelton turbine or a pressure exchanger system (PES), a slightly lower EC, about 27.5 kWh/m 3, was obtained. However, when thermal energy was available in the plant or the stream was already at the operating temperature of the MCr unit, the EC decreased by an order of magnitude; i.e. to 2.05 kWh/m3 without energy recovery and to about 1.6 kWh/m 3 with heat recovery. The authors [22] estimated also the exergetic efficiency considering different hybrid systems with and without energy recovery device (ERD). The efficiency was found to be higher for the hybrid systems without MCr and MD than for the hybrid systems including MCr and MD. However, for all considered cases, the efficiencies were far better than 1.12–10.4%, which was for MSF. Al-Obaidani et al. [30] performed exergy calculations for a 24,000 m 3/day DCMD desalination plant operated with and without heat recovery system to reuse the heat from the brine to preheat


the feed aqueous seawater solution with a heat recovery efficiency of 80%. The heat input was 45,036 kW and 39,690 kW for the plant with and without a heat recovery system. The calculated exergy efficiency was 28.3% and 25.6% for the desalination plant with and without heat recovery, respectively, while the net exergy change between inlet and outlet was 71 and 353 kW with and without heat recovery, respectively. The EC was 39.7 kWh/m 3 and 45 kWh/m 3 for the plant operated with and without heat recovery, respectively. For MD technology, the EC is much higher for small laboratory MD systems compared to larger pilot plants with greater membrane areas. When heat recovery systems are used the EC could be reduced considerably. Until now in most of the studies rough estimations are performed and most of them are only theoretical. Analysis based on realistic energy consumption for MD pilot plants is required in the future. For sake of comparison, Table 3 shows the EC of other separation processes used in desalination such as RO, ED, MSF, solar still, MED and VC. RO technology exhibits the lowest EC values than the other processes including MD [63]. At present, RO is the most energy-efficient technology for seawater desalination. This is attributed mainly to its continual technological advances and improvements including the decrease of the membrane module costs, the development of energy‐efficient membrane modules and pumps as well as to more efficient energy recovery devices. These justify the worldwide use of RO for industrial seawater desalination. 3. Dispersion in the MD costs evaluations Few researchers have reported water production costs (WPC) of MD [8,10–12,17,19,20,22,23,29,30,32]. This is because of various reasons: (i) the costs of MD process equipment and its components

Table 3 Specific energy consumption, EC, of different separation processes used in desalination. EC (kWh/m3)

Observations

Ref.

Reverse osmosis (RO) 1–2 RO, brackish water. 1.3 RO, brackish water with photovoltaic panels and a thermal collector, without energy recovery, water production: 0.2 m3/day. 3 RO, brackish water. 2.5–4.5 RO, water production: 105,000 m3/day. 2–4 RO, seawater. 3.9–5.6 RO, seawater. 5 RO, seawater. 5–6 RO, seawater with energy recovery. 7.5 RO, seawater. 17 RO, seawater, electric energy. 17 RO, seawater. 22.8 RO powered with photovoltaic panels.

[26] [20] [58,60] [59] [61] [26] [54] [26] [53,62] [17]

Multistage flash (MSF) 26.4 MSF without energy recovery device 41 MSF 60–80 MSF 63.9 MSF 83 MSF, seawater.

[14] [20] [26] [54] [53,62]

[58,60] [52]

Vapor compression (VC) 7–19 Single effect evaporator, mechanical vapor compression (VC) desalination system. 7.3–12.5 VC desalination system. 10.4–11.2 VC desalination plant, water production: 500 m3/day.

[56] [57]

Simple effect distillation 640 Solar still (SS).

[26]

Multi-effect distillation (MED) 416.7 Multi-effect solar still (MESS). ≈30 MED thermal energy.

[17] [26]

[55]

95

such as MD modules are not yet known precisely, (ii) the technology is still under test and not fully applied in industrial scale, (iii) there is not an unified standard method for computing and reporting the costs in MD, (iv) there is a wide dispersion in the EC, and (v) some costs were neglected such as land cost, pretreatment cost, etc. It may be expected that the lower is the EC the more economical is the MD process. However, the total cost for MD involves different strongly interrelated factors that depend upon: i) Source water (type of feed water to be treated, seawater or brackish water, salinity and quality of feed water): Low salinity of feed leads to high distilled water production and conversion rates. Consequently the plant can operate with low EC, less antiscalent chemicals and less cleaning cycles. ii) Energy sources: Both electrical and thermal are used in MD. Available and inexpensive sources for heating and low cost electric power affect considerably the WPC. iii) Capital: MD process equipment, membrane modules, installation and building, control instrumentation, land or rental, auxiliary equipment, renewable energy conversion and storage, etc. iv) Plant life, amortization or fixed charges: Increase in plant lifetime reduces the capital cost (CC). v) Operation: pretreatment, post-treatment, brine or concentrate disposal, etc. vi) Maintenance: cleaning, membrane replacement, staff or labor, etc. vii) Plant size and capacity: Increase of the MD plant capacity may decrease the WPC up to a minimum. Optimum plant capacity should be determined for a specific MD plant type. Fig. 1 shows a summary of the cost model breakdown of MD process. As can be seen the WPC is divided in 3 groups namely, direct capital cost (DCC), indirect capital cost (ICC) and annual operating cost (AOC). Actually, the MD process equipment is the most cost items that depend on the plant capacity as well as on the MD configuration. At present the MD plant costs are very expensive due to the lack of competitiveness. The land cost and the building cost depend on the site specific. The installation cost and the control instrumentation cost is assumed to be 25% the cost to purchase equipment. Furthermore, land or rental cost should be included. However, in MD most of authors considered zero land cost [29]. The pumping cost is related to the pressure drop in the MD membrane module and the cost of MD membrane varies widely because the membrane is still under development [6,7]. The used membranes were prepared for other technologies, microfiltration (MF) or ultrafiltration (UF). For instance, different costs were assumed for the MD membranes. Macedonio et al. [22] and Al-Obaidani et al. [30] assumed membrane cost as $90/m 2. Drioli et al. [12] considered a higher membrane cost of $116/m 2, whereas Banat and Jwaied [29] considered a cheaper cost for the membrane, $36/m 2, which is 33.3% the total cost of the membrane module (i.e. membrane cost + membrane assembly cost). However, this price may change depending on the type of membrane (pore size, material, etc.), company and the quantity purchased (plant capacity). Banat and Jwaied [29] reported a membrane module cost equal to 13.6% of the total cost of a solar-driven AGMD compact plant ($7965) with a capacity of 100 l/day (10 m 2 membrane area) and a lower membrane module cost 8.2% of the total cost for a solar-driven large plant ($52,380) with a capacity of 500 l/day (40 m 2 membrane area). All items indicated in the ICC are expressed as percentage of the DCC that may vary in the range 5–15%. For example, for insurance cost 5% of DCC can be considered, whereas the cost of construction overhead is equal to 15% of direct material and labor cost and then adjusted for the size of the plant (capital cost, CC). For project contingency 10% of DCC can be assigned. The costs of administration and legal fees are equal to 10% of CC. For MD technology, the ICC is considered 10% of the DCC [22].

96


Direct capital cost (DCC) - MD Process equipment (membrane modules, heat exchangers, pumps). - Auxiliary Equipment (open intakes, pipe lines, valves, transmission piping, storage tanks, generators and transformers, electric wiring, brine disposal line). - Systems pre-treatment - Systems post-treatment. - Installation and building (workshop, control room, laboratory, water analysis). - Renewable energy conversion and storage - Land cost

Water production cost (WPC)

Indirect capital cost (ICC) - Administration and legal fees - Insurance - Construction overhead - Labor burden - Field supervision - Temporary facilities - Small tools and miscellaneous - Contingency

Annual operating cost (AOC) - Grid energy - Labor - Membrane replacement - Operation and maintenance (O&M) - Brine disposal - Consumables and chemicals for pretreatment and post-treatment: (Antiscalent, cleaning agents, etc.). - Amortization or fixed charges. Fig. 1. Cost elements needed to determine the WPC of MD process.

The annual operating costs (AOC) or running costs are the total yearly costs of owning and operating a desalting plant including amortization or fixed charges, operation and maintenance (O&M) costs, membrane replacement costs (CMR), consumables, etc. The labor cost is site specific and depends whether the plant is government or private. etc. The CMR is a function of the produced water quality and will vary depending on the type of membrane and module, MD process configuration, feed salinity, use of pretreatment system, etc. Therefore, the membrane replacement time per year will vary. Nowadays, one can only speculate on the membrane lifetime. Zuo et al. [36] considered a membrane life of 5 years. Banat and Jwaied [29] estimated the CMR to be 20% of the membrane module cost. For example, for a solar-driven AGMD compact plant with 100 l/day capacity the calculated CMR value is $216/year, whereas for a larger capacity of 500 l/day it is $864/year. The yearly rate for the maintenance cost can be considered equal to 2% of the CC. The cost of chemicals depends on the availability of nearby manufacturing industries and feed type. The pretreatment cost depends on the water source and type of the feed aqueous solution to be treated. In some MD studies, authors assumed zero pretreatment cost [29,36]. The amortization or fixed charges (Cfixed) accounts for the annual interest payments of the DCC and ICC. It can be calculated by the amortization factor a, given by: n

a¼

ið1 þ iÞ ð1 þ iÞn −1

ð13Þ

where i is the annual interest rate (%) and n is the year or lifetime of the plant. Banat and Jwaid [29] considered i = 0.05 (5%), n = 20 and

found an amortization factor of 0.080243/year. These values assigned to i and n were also considered by Zuo et al. [36]. However, today there is a lack of experience in the MD desalination industry to assign the amortization life of 20 years. For the interest rate, a range of 3–8% may be considered. The annual fixed charges can be estimated from the CC using the following equation: C fixed ¼ aCC

ð14Þ

For the solar-driven compact AGMD plant, the calculated Cfixed is $639/year [29]. The O&M costs (CO&M) accounts for the annual payments for the operation and maintenance of the plant, staff cost, spare costs etc. In MD, the annual CO&M was estimated to be 20% of the plant annual payment [29]. C O&M ¼ 0:2C fixed

ð15Þ

For a solar-driven AGMD compact plant with 100 l/day capacity the calculated AO&M is $128/year and for a higher AGMD plant capacity of 500 l/day it is $841/year. The plant availability (f) refers to the working time of the plant and can be assumed to be 90% per year [22,29,36]. The total annual cost (Ctotal) can be calculated as: C total ¼ C fixed þ C O&M

ð16Þ


The WPC can be determined using the following expression: WPC ¼

C total f M365

ð17Þ

where f is the plant availability and M the plant capacity. For example, the obtained Ctotal for a solar-driven AGMD compact plant with 100 l/day capacity is $983/year and its WPC is $29.9/m 3. However, for a larger solar-driven AGMD plant with 500 l/day capacity Ctotal is higher ($5908/year) and the WPC is $36/m 3. It is worth quoting that other AOC should be considered such as annual labor cost (Clabor) and annual brine disposal (Cbrine). The annual labor cost (Clabor) can be calculated as follows: C labor ¼ gMf 365

ð18Þ

where g is the specific cost of operating labor assumed as $0.05/m3 [22]. The annual brine disposal (Cbrine) is expressed as: C brine ¼ bMf 365

ð19Þ

where b is the specific cost of brine disposal assumed as $0.0015/m 3 [30]. The annual electric power cost (Celectric) can be calculated using the following expression [22]: C electric ¼ cwMf 365

ð20Þ

where c is the electric cost about $0.09/kWh and w is the specific consumption of electric power (kWh/m 3) [22]. Gilron et al. [24] and Zuo et al. [36] considered an electric cost of $0.06/kWh. Recently, much more detailed cost analysis studies in MD have been carried out by Macedonio et al. [22], Banat and Jwaied [29] and Al-Obaidani et al. [30]. Data and equations to be used for economic calculations as well as the assumptions involved were reported [22,29,30]. In order to figure out the way to lower the cost of desalination by MD, the factors that contribute appreciably on the total cost must be first understood. Nowadays, the major cost elements for MD desalination plants are the CC [7,29]. Therefore, improvements must be made first in CC to lower the WPC of MD. Table 4 summarizes the WPC estimated for various MD systems. As can be seen the WPC values vary in 4 orders of magnitude from $0.3/m 3 to $130/m 3. Distinct WPC values were claimed and reported

97

depending on the type and size of the MD system, feed processed water, energy source, energy recovery systems, cost of energy, economic analysis procedure, etc. About 25 years ago, cost analysis of a DCMD plant with heat recovery producing 5000 kg/h (44,000 t/year) fresh water was performed by Fane et al. [9]. Hollow fiber membrane modules with a total surface area of 800 m2 and three heat exchangers and two pumps were considered in their calculations. Scaled to 1987 $US, the calculated installed capital cost was $720,000, the operating cost was $3.3/t and the total cost was $4.9/t. According to their calculation, the WPC decreased considerably with the increase of the applied feed temperature and at 90 °C feed temperature the WPC could be as low as $2/m3. Moreover, the costs for a small solar-driven MD pilot plant of production capacity of 50 kg/h (500 kg/day, daily production over 10 h) were estimated. They found that a capital cost would be in the range $10,000 to $15,000 with a WPC of $10/m 3 to $15/m3. About 20 years ago Hogan et al. [10] studied the feasibility of a solar‐powered MD pilot plant with heat recovery for the supply of domestic drinking water in the arid/rural regions of Australia. They optimized solar collector area, membrane area and heat recovery to achieve low CC and high water production. It was found that the CC was very sensitive to the heat recovery factor and for a minimum CC the heat recovery factor should be between 60 and 80%. For a production capacity of 50 kg/day the optimum configuration was a solar collector area of around 3 m 2, a membrane area of 1.8 m 2 and a total heat exchanger area of 0.7 m 2 with a CC of $3500 (Australian in 1991). Sarti et al. [11] estimated the cost of benzene removal from wastewater containing 1000 ppm of benzene by VMD plant with 5 stages. In each stage the feed aqueous solution was the retentate of the previous stage. The CC of the plant was $247,000 designed for 99% benzene removal with heat recovery of the retentate to preheat the liquid wastewater. Capital depreciation, labor cost, module replacement and energy consumption were taken into account to estimate the treatment cost per unit volume of wastewater. The labor cost was 10% of the CC per year, the membrane cost was $450/m 2 in module, the assumed membrane life was 3 years, the depreciation was 15% of CC per year, the operation time was 7200 h per year, the pump efficiency was 0.8, the electricity cost was 0.085 $/kWh, the steam at low pressure was 0.013 $/kg and the cooling water cost was $15/m3. The estimated WPC was $4.04/m 3 [11]. In 1999, Drioli et al. [12] performed cost analysis of RO/MD hybrid system for water desalination. MD was proposed to treat RO brine with a concentration of 75 g/l at a temperature of 35 °C in order to enhance both efficiency and water recovery factor in seawater

Table 4 Estimated water production cost, WPC, of different MD systems. WPC ($/m3)

Year

Observations

Ref.

10–15 4.04 1.25 3.2 1.2 130 0.26–0.54

1987 1993 1999 2003 2003 2005 2006

[9] [11] [12] [15] [15] [17] [20]

0.92 0.74 0.55 0.71 0.51 ≈1 1.23 1.17 0.64 15–29.9 18–36

2007 2007 2007 2007 2007 2007 2008 2008 2008 2008 2008

Solar-driven DCMD plant, water production: 500 kg/day. VMD plant for benzene removal. MD hybrid system (RO&MD), WPC: $1.25/m3 for only RO, WPC: $1.32/m3 for only MD. VMD single-pass flow, permeate flux: 0.7 l/m2.h. VMD discontinuous flow, permeate flux: 0.5–0.7 l/m2.h. AGMD, use of sensible heat of geothermal, water production: 17 l/day. Memstill® units, water production: 105,000 m3/day, heat supply generated by cogeneration of heat and electricity, fuel fired or by a waste heat source. MD hybrid system (NF/RO/VMD), water production: 76.2 m3/day. MD hybrid system (MF/NF/MCr/RO/MD) without energy recovery device, thermal energy available in the plant or steam. MD hybrid system (MF/NF/MCr/RO/MD) with energy recovery device. MD hybrid system (MF/NF/MCr/RO/MD) without energy recovery device, thermal energy available in the plant or steam. MD hybrid system (MF/NF/MCr/RO/MD) with energy recovery device, thermal energy available in the plant or steam. AGMD, permeate flux: 5.2 l/m2.h. MD with heat recovery, water production: 24,000 m3/day. MD without heat recovery, water production: 24,000 m3/day. MD with low-grade heat energy source. Compact AGMD solar-driven plant, water production: 0.1 m3/day. Large AGMD solar-driven plant, water production: 0.5 m3/day.

[23] [22] [22] [22] [22] [21] [30] [30] [30] [29] [29]

98


desalination (45 g/l at 25 °C). It was found that RO/MD hybrid process produced more than twice as much water as RO process alone at the same water cost, and the MD process alone produced as much water as the RO/MD hybrid process but at a water cost about 5% higher (i.e. $1.25/m 3 for RO alone with a production 0.391 time the feed flow rate, $1.32/m 3 for MD alone with a production 0.856 time the feed flow rate and $1.25/m 3 for RO/MD with a production 0.856 time the feed flow rate). Specific and detailed cost analysis was not indicated by Drioli et al. [12]. During last 10 years, some studies have been conducted on MD economics [17,20,22–24,29,30,32,36,64]. Bouguecha et al. [17] estimated the annual cost of 17 l/day AGMD process using sensible heat of geothermal water resource. The CC of the plants as well as the costs of O&M was taken into consideration. The O&M cost (CO&M) was estimated as 1/6 of annual cost. The obtained WPC was $130/m3 (i.e. $110/m3 CC and $20/m3 CO&M). Again, detailed calculations were not reported by Bouguecha et al. [17]. The WPC of Memstill® technology was reported by Meindersma et al. [20] and the results were compared to those of RO for seawater desalination capacity of 105,000 m 3/day. The heat supply to the AGMD process was made by cogeneration of heat and electricity, fuel combustion or by a waste heat source. Compared to RO, lower WPC (i.e. $0.26/m 3) for Memstill® was reported when cheap waste heat of $0.1/GJ was considered. Future prospect of WPC for different desalination processes including MD was made [20]. It was expected for next coming years, the WPC of MD would be lower than that of RO, the water costs by MSF and MED would not be below $1/m 3, and therefore only RO would be able to compete with MD in the coming future. The process promises to decrease desalination costs to well below $0.50/m 3 if low‐grade waste steam or heat is used. However, Memstill® technology is still under test and should be proven on a larger scale before confirming the projected decrease of desalination costs by MD [19,20]. Wang et al. [32] integrated a recirculating cooling water (RCW) to a DCMD process for water production. Total investment and O&M costs were estimated for the MD hybrid system with a capacity of 50 m 3/h and compared to those of a hybrid unit, UF/RO integrated to RCW with a water recovery rate of 70% and the same capacity. Lower CC (280,000 US$) and CO&M ($0.139/m 3) were obtained for the MD hybrid plant compared to those of UF/RO hybrid plant (320,000 US$, $0.504/m 3). The CO&M included the cost of electricity ($0.067/kwh), chemicals and labors ($400/month.person) and the lifetime of all facilities was 15 years, whereas that of the membrane was 3 years. El-Zanati and El-Khatib [23] proposed a hybrid system composed of NF, RO and VMD for water desalination. NF and RO units were combined together to form the feed solution of the MD unit, while the produced water by RO and MD processes was mixed giving the last product. Simulations and cost analysis have been carried out for a water production of 76.2 m 3/day. It was found that a WPC would be about $0.92/m 3, which was claimed to be competitive to potable water produced by seawater RO systems. Again, in this case rigorous cost analysis was not reported. Macedonio et al. [22] also performed the economic evaluation of seven systems (RO alone and different RO hybrid systems) for seawater desalination. One of the proposed system integrated MF, NF, MCr, RO and MD in one hybrid system. MCr was applied on NF retentate and MD was applied on RO retentate. Different possibilities have been considered in terms of heat recovery and available thermal energy in the plant. It was found that the WPC decreased from $0.74/m 3 to $0.55/m 3 when the thermal energy is already available in the MD hybrid plant. The cost was lower with both heat recovery and available thermal energy in the plant, whereas higher cost was with neither energy recovery nor available thermal energy in the plant. For the MD hybrid system integrating MCr, when the gain for salts sale was taken into consideration, negative WPC ($ − 0.36/m 3

to $ − 0.13/m 3) were obtained [22]. The authors claimed that the gain for the salts sale covered more than entirely the cost of WPC and therefore advised the integration of MCr in the desalination plant. Al-Obaidani et al. [30] made calculations for a DCMD plant with a capacity of 24,000 m 3/day, with and without heat recovery. A sensitivity test for different variables of DCMD on the process economics was also performed in order to identify the most sensitive parameters on the WPC and to find the optimal operational conditions. One of such studies was the effect of the temperature difference between the feed input and the permeate output at the membrane module(s). It is well known that an increase in the temperature difference (i.e. driving force in DCMD) enhances the permeate flux and therefore the membrane area may be reduced and consequently the CC is decreased. However, a great temperature difference requires a large amount of heat input on the feed aqueous solution, which in turn increases the CO&M and the WPC. Considering these two effects of the temperature difference, an optimum point may be expected. The minimum WPC obtained by Al-Obaidani et al. [30] was $1.23/ m 3 for the feed inlet temperature of 55 °C (permeate outlet temperature of 25 °C) for DCMD without heat recovery. With heat recovery the minimum WPC was $1.17/m 3 when the feed inlet temperature was 60 °C and the permeate outlet temperature was 30 °C [30]. It was reported that the membrane cost contributed about 50% of the CC and 30% of the CO&M in the DCMD plant [30]. The effect of the membrane cost on the WPC was studied and it was observed that when the membrane cost was increased by 10%, the WPC increased 3.9% (with heat recovery) and 3.7% (without heat recovery). In addition, the WPC was found to be sensitive to the steam cost. When the price of the steam increased 10%, the unit cost of water increased by 4.4% for the MD operation with heat recovery and 5% for the MD operation without heat recovery. Al-Obaidani et al. [30] concluded that the WPC became $1.23/m 3 and $1.17/m 3, respectively, for the MD operation without and with heat recovery at the optimum temperature difference. Finally, it was reported that the WPC would fall down to $0.64/m 3 if waste heat was used. Banat and Jwaied [29] made economic evaluation of two solardriven AGMD plants using spiral wound membrane module(s) with heat recovery (Fraunhofer ISE, Germany). One is a compact unit installed in the northern part of Jordan (Irbid) and has been operated with brackish water since September 2005 and the other is a large unit installed in the southern part of Jordan (Aqaba port) and has been operated with untreated seawater since February 2006 [25,26]. Both units consist of flat plate thermal collectors and photovoltaic (PV) panels and data acquisition systems. The technical characteristics of each desalination plant together with the CC and CO&M were reported in [29]. The EC of this type of plants was discussed in the previous section. The obtained WPCs were $29.9/m 3 and $36/m 3 for the compact and large AGMD plants, respectively. It was reported that this cost was for very pure water and the production cost of drinkable water with a salt content of 500 ppm would decrease the WPC to one half, $15/m 3 for the compact unit and $18/m 3 for the large unit [29]. Banat and Jwaied [29] have also calculated the WPC when there is no interest payment required. Then, the cost of drinkable water became even lower, $10.2/m 3 and $11.9/m 3 for the compact and large system, respectively. Moreover, the influence of membrane and plant lifetime on WPC was simulated observing a significant reduction of WPC with increasing lifetime of the membrane and plant. For example, for 1 year membrane lifetime the WPC was found to be close to $30/m 3 for both compact and large AGMD plants; however for 7 years the WPC was $14/m 3 and $17/m 3 for the compact and large plants, respectively. Therefore, increasing the reliability of the MD technology and MD membrane and plant lifetime could reduce the cost significantly. Kullab and Martín [64] reported simulated WPC for AGMD plants ranging from $1.3/m3 to $1.5/m3 for a total membrane area of 2392 m2,


and from $0.6/m3 to $0.7/m3 for a lower membrane area of 1141 m2 with the same water production rate 10 m3/h. Gilron et al. [24] optimized the WPC with the number of stages of cross-flow DCMD modules for an individual cascade and found it to be between 4 and 6 with a WPC of $1.6/m 3 assuming a steam cost of $4/mt. However, in the case of waste steam with $1.76/mt, the WPC was found to be lower than $0.9/m 3 with a minimum for a number of stages between 4 and 6 of about $0.75/m 3, and the number of stages was reduced from 6 to 2 as the energy price was decreased from $4/mt to $0.73/mt. Based on several assumptions, the theoretical relationships between the WPC and the membrane area as well as the operation parameters of a cross-flow DCMD system such as the feed and permeate flow rates were studied by Zuo et al. [36]. To evaluate the WPC only 5 main components were considered (i.e. membrane module, heat exchanger, electricity consumption for pumping, waste steam and installation). The CO&M was not considered. The estimated WPC values were found to be less than $1.8/m 3, which was within the WPC range reported by Gilron et al. [24]. For a membrane area around 4 m 2 a minimum WPC of $1.1/m3 was estimated. Below this membrane area considerable increases of water production and the GOR were observed with increasing membrane area leading to a significant drop in the WPC. However, above 4 m 2 membrane area, less effect was detected on the water production and the GOR, and therefore the increase in the membrane cost became the dominant factor causing an increase of the WPC. In addition, with increasing feed and permeate flow rates the estimated WPC was reduced gradually achieving asymptotic values and indicating that it was favorable to operate the DCMD system at moderate flow rates because significant electric energy consumption was required for high flow rates. Table 5 summarizes the WPC of other desalination processes applied using different energy sources. It is worth quoting that the WPC of desalination has fallen considerably since the past 50 years. The WPC of desalting seawater has been reduced to about $1.0/m 3 or less, while the WPC of desalting brackish water has decreased to about $0.6/m 3 [1]. In fact, the WPC of MD technology is still higher than that exhibited by other desalination processes. Among them, the dominant competing process is RO that exhibits a WPC between $2.37/m 3 for a plant with capacity of 0.3 million gallons per day to $0.55/m 3 for 30 m gallons/day capacity system [46]. Table 5 clearly shows that RO exhibits a significant economic advantage for treating brackish waters. For desalination of seawater, RO has an economic advantage over MSF. The average WPC has declined considerably over time from $5.0/m 3 in 1970 to less than $1.0/m3 [1]. The WPC for seawater desalination is still above $1/m3 and the WPC for desalting brackish water is below $0.6/m 3 [1]. This is due to the considerable improvements of RO technology during the past 40 years. For MSF technology, the WPC also has been reduced substantially from an average of about $9.0/m3 in 1960 to about $1.0/m 3 [1]. For the VC process, the WPC has decreased from $5.0/m3 in 1970 to about $1.0/m3 [1], whereas the WPC of desalination by MED process has fallen from $10.0/m 3 in 1950s to about $1.0/m 3 [1]. 4. Conclusions An overview of the studies performed on the gained output ratio (GOR), specific energy consumption (EC) and the water production costs (WPC) of different MD systems are presented together with comparisons to other desalination processes. So far all the published studies on the economics, energy analysis and costs evaluations are about MD desalination of seawater and/or brackish waters. In fact, very few studies were published on these issues and a wide dispersion in the EC and the WPC of MD technology has been observed. The EC varies from about 1 to 9000 kWh/m 3 (3 orders of magnitude dispersion) or even higher while the WPC varies from $0.3/m 3 to $130/m 3 (4 orders of magnitude fluctuations). Distinct WPC values

99

Table 5 Estimated water production cost, WPC, of different separation processes used in desalination. WPC ($/m3)

Observations

Reverse osmosis (RO) 0.11 RO, electric, water production: 39,000 m3/day. 0.45–0.63 RO, water production: 105,000 m3/day. 2.7 RO, solar photovoltaic panels, water production: 500 m3/day. 3.73 RO, solar photovoltaic panels, water production: 1 m3/day. 12.05 RO, solar photovoltaic panels, water production: 1 m3/day. 80 RO, solar photovoltaic panels. 0.25–0.28 RO, brackish water. 0.20–0.35 RO, brackish water. 0.45–0.92 RO, seawater. 0.55–0.80 RO, seawater. 0.7–0.8 RO, seawater. 1.25 RO, seawater. 1.51–6.56 RO, seawater. Multistage 1.10–1.50 1.31–5.36 ≈1.4 2.02 2.84

flash (MSF) MSF MSF MSF MSF, natural gas, water production: 20,000 m3/day. MSF, solar, water production: 1 m3/day.

Reference

[68] [20] [67] [66] [65] [17] [78] [58,74] [58,74] [78] [59] [12] [53,58]

[58,74] [53,58] [30] [69] [65]

Vapor compression (VC) 0.46 VC 0.87–0.92 VC 0.99–1.21 VC

[75] [58,74] [77]

Simple effect distillation 12 Solar still (SS), water production: 1 m3/day.

[70]

Multi-effect 0.46–85 0.89 ≈1 1.17 2 7–10 50

[74,58] [73] [30] [76] [71] [72] [17]

distillation (MED) MED MED, solar, water production: 20,000 m3/day. MED MED MED, solar, water production: 72 m3/d. MED, solar, water production: 85 m3/d. Multi-effect solar still (MESS)

were claimed and reported depending on the type and size of the MD system, feed processed water, energy source, energy recovery systems, cost of energy, economic analysis procedure, etc. In some reported MD papers authors dedicated only a small section to the EC and the WPC without specifying the followed calculations of the costs and energy analysis. In fact, there is no agreed on standard for calculations of WPC. The EC is much higher for small laboratory MD systems compared to larger pilot plants with greater membrane areas. Until now in most of the published studies rough estimations of EC and WPC are performed and most of them are only theoretical. Analysis based on realistic EC and WPC of MD pilot plants is required. The observed dispersions in EC and WPC may be due to various reasons. The MD is not yet fully applied for commercial scale and the capital investment costs that include membranes and modules are fluctuating. Other cost-related information such as pretreatments, optimum flow conditions, long-term MD performance, fouling, and membrane life are not yet available at a satisfying level. The highly concentrated brine discharged by the process is also of ecological concern and its related cost should be taken into consideration. Optimized MD plants must be designed and developed first and then rigorous cost and economic analysis for comparison will become possible. More intensive and focused MD research efforts in this field are needed, not only theoretically but also experimentally looking forward to decrease the EC and WPC of MD technology. The economic penalty of MD comes mainly from the high initial capital investment, especially the MD membrane modules. At present, the MD membranes and modules are expensive.

100


Very few membrane modules and pilot plants are being commercialized by some companies such as Scarab Development AB, Memsys and Fraunhofer ISE. Much larger MD plants capacities should be constructed and a unified standard method for economic analysis procedure should be followed to determine the WPC. Various factors should be considered in order to estimate adequately the WPC of an MD installation, including direct and indirect capital costs as well as annual operating and maintenance costs. For the benefit of MD process, one should be cautious when reporting simulated, non-realistic and non-contrasted WPC. For MD systems more investigations on economics of the process should be conducted. In the next 10–20 years it is expected to see a massive increase in capacity and production of water by MD technology. Therefore, MD desalination capacity is projected to grow with a decrease in EC and WPC. These reductions are attributed to the continual technological improvements in the field of MD, including higher permeability membranes, efficient modules and energy recovery devices, multistaged MD effect, hybrid installations, use of renewable energy systems and cheaper waste heat, brine management technologies, etc. Symbols a amortization factor A membrane area AOC annual operating cost b specific cost of brine disposal c electric cost Cbrine annual brine disposal CC capital cost Celectric annual electric power cost Cfixed annual fixed charges Clabor labor cost CMR membrane replacement cost CO&M O&M costs cp specific heat Ctotal total annual cost DCC direct capital cost EC specific energy consumption Ee electrical energy Ein global energy input Et thermal energy EX exergy of a flow stream EXT temperature exergy term EXP pressure exergy term EXC concentration exergy term Exout exergy input Exin exergy output f plant availability g specific cost of operating labor GOR gained output ratio HR heat recovery factor i annual interest rate ICC indirect capital cost Jw permeate flux M plant capacity _f m feed flow rate _ m mass flow rate n year or lifetime of the plant. ns solvent concentration (mol/kg solution) P pressure QC heat flux transferred by conduction Qf heat transfer in the feed channel QHE maximum heat recoverable in the main heat recovery exchanger Qm total heat flux transferred through the membrane QMD heat transferred in the membrane module

QV R T Tf,in Tf,out w WPC xs

heat flue associated to permeate flux gas constant absolute temperature temperature of the feed solution at the inlet of the membrane module temperature of the feed solution at the outlet of the membrane module specific consumption of electric power water production cost solvent mass fraction

Greek letters ΔHv,w enthalpy of evaporation of water ΔTHE temperature difference across the heat exchanger ΔTMD temperature difference across the membrane ΔTMD,module axial temperature drop along the MD module εE energy efficiency εT thermal efficiency ρ density ξ exergetic efficiency

Abbreviations AGMD air gap membrane distillation DCMD direct contact membrane distillation ED electrodialysis ERD energy recovery device MCr membrane crystallization MD membrane distillation ME multiple effect evaporation MED multi-effect distillation MF microfiltration MSF multistage flash NF nanofiltration O&M operation and maintenance PV photovoltaic panels RO reverse osmosis SS solar still VC vapor compression VMD vacuum membrane distillation

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