A novel Forward osmosis membrane pretreatment of seawater for thermal desalination processes

A novel Forward osmosis membrane pretreatment of seawater for thermal desalination processes

Desalination 326 (2013) 19–29 Contents lists available at ScienceDirect Desalination journal homepage: www.elsevier.com/locate/desal A novel Forwar...

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Desalination 326 (2013) 19–29

Contents lists available at ScienceDirect

Desalination journal homepage: www.elsevier.com/locate/desal

A novel Forward osmosis membrane pretreatment of seawater for thermal desalination processes Ali Altaee a,⁎, Abdelnasser Mabrouk b, Karim Bourouni c a b c

Faculty of Engineering and Physical Sciences, University of West of Scotland, Paisley, PA1 2BE, UK Engineering Science Department, Faculty of Petroleum & Mining Engineering, Suez University, Suez, Egypt Maître Assistant à l'ENIT, Département de Génie Industriel, 1012 Tunis, Tunisia

H I G H L I G H T S • • • • •

FO pretreatment of SW to thermal desalination was suggested. FO–MSF/MED design was optimized. Model for FO process was developed. CaCO3 scale model was developed. FO model coupled with VDS program to calculate thermal plant performance.

a r t i c l e

i n f o

Article history: Received 30 May 2013 Received in revised form 9 July 2013 Accepted 10 July 2013 Available online 3 August 2013 Keywords: FO membrane softening Seawater pretreatment Thermal desalination Scale removal

a b s t r a c t The present work introduced a novel conceptual design of integrating Forward Osmosis (FO) membrane with the Multi Stage Flashing (MSF) or Multi Effect Distillation (MED) thermal desalination processes. A simple mathematical model was developed here to estimate the performance of the FO membrane system. A previously developed program, VDS, for estimating the performance of thermal processes was updated to include the FO system. The verified VDS program [1] was applied to simulate the performance of the FO–MSF/MED hybrid system at different recovery rates varied from 16% to 32%. Brine reject from the thermal desalination processes was recycled and used as a draw solution to reduce the cost of FO membrane pretreatment. Seawater was used as the donor solution in the FO membrane. The simulation results showed that the FO pretreatment, successfully, reduced the concentration of multivalent ions in the feed solution to the MSF and MED. It was found that the concentrations of Ca2+, Mg2+, and SO24 − ions, which are responsible for scale problem in MSF, decreased with increasing the recovery rate of FO membrane. In case of FO–MED hybrid system, the thickness of the CaCO3 scale layer was calculated at different FO recovery rates. The estimated thickness of CaCO3 scale layer was 74 μm, 43 μm, and 39 μm for 0%, 20%, and 32% FO recovery rate respectively. It was also found that the thickness of CaCO3 scale layer decreased in the direction from effect 1 to effect 6 due to temperature drop. Finally, the study demonstrated the feasible application of FO membrane in the pretreatment of seawater to reduce the concentration of multivalent ions which are responsible for the scale problem in the thermal desalination processes. Crown Copyright © 2013 Published by Elsevier B.V. All rights reserved.

1. Introduction Thermal desalination processes have been used in seawater desalination in the region of the Middle East since the early sixties of the last century [1]. The current worldwide market share of the thermal desalination processes is about 35% while RO membrane technology represents 61% [2]. However, the market share in Gulf Cooperation Council (GCC) countries showed that thermal technologies represent 70% and RO membrane technology covers the remaining cake.

⁎ Corresponding author. Tel.: +44 7986517994. E-mail address: [email protected] (A. Altaee).

This situation reflects the thermal process maturity for large capacity production with high purity. The harsh gulf seawater property (high temperature, high salinity, high impurity and sometimes red tide) increases the cost of RO membrane replacement and pretreatment cost [2]. Multi stage flushing (MSF) and Multi Effect Distillation (MED) are the major thermal technologies for seawater desalination and fresh water supply in a number of countries in the Middle East. Despite the high efficiency of the thermal processes, they suffer from an essential drawback represented in the deposition and accumulation of scale materials on the surfaces of heat exchangers [3–5]. Normally, none-alkaline scale is deposited on the internal pipe surfaces in the MSF while alkaline scale deposit was reported on the external surface of heat exchangers in

0011-9164/$ – see front matter. Crown Copyright © 2013 Published by Elsevier B.V. All rights reserved. http://dx.doi.org/10.1016/j.desal.2013.07.008

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the MED [4–6]. Scale deposition in thermal processes is renown of reducing the heat transfer efficiency in the evaporator and hence the process efficiency [4,6]. To avoid this problem, antiscalants are normally dosed into the feed water to the thermal processes. In real life scale deposition can't be completely avoided even when antiscalants is used and hence mechanical and chemical cleaning is often applied to clean the fouled parts. Furthermore, scale deposition restricts the Top Brine Temperature (TBT) increase beyond the precipitation point of the inversely soluble metal ions. Accordingly, the TBT's of MSF and MED cannot exceed 112 °C and 65 °C respectively [4,7,8]. In the nineties of past century, Ata and co-workers suggested the application of Nanofiltration membranes (NF) for the removal of multivalent ions from feed water to the thermal desalination processes [3,4]. NF membranes have the ability to remove divalent ions from seawater which are responsible for scale deposition in thermal units such as Ca2+, Mg2+, and SO2− 4 . Pilot and bench scale experiments showed the feasibility of NF pretreatment in the removal of scale ions from seawater and hence it was possible for the MSF unit to operate at TBT's over 112 °C [3,4]. At higher TBT's, the Gain Output Ratio (GOR) of MSF was increased over 8. Further research on NF pretreatment proposed a partial dilution of the feed solution to the thermal desalination processes with NF permeate in order to reduce the pretreatment cost. The pilot plant tests demonstrated more than 75% removal of scale ions from the seawater feed to the thermal desalination processes. After scrutiny, however, it was found that the cost of NF was uneconomical even at NF recovery rate more than 65% [7]. Technical problems were also reported with the NF pretreatment of seawater in Sharjah pilot plant test, UAE [7]. Therefore, the divalent ions removal from seawater by NF would only be possible if the cost of membrane pretreatment is reduced. The techno-economic analysis of integrating NF pretreatment for the existing multi stage flash-brine recirculation (MSF-BR) and newly developed MSF-DM configurations was evaluated [8]. The cost analysis showed the unit product cost was 5 % higher than that conventional MSF (at 110 °C) due to the additional capital cost of NF system. When NF system was integrated in a new desalination plant configuration NF-MSF-Deaeration and Brine Recycle (NFMSF-DM) at TBT = 130 °C, the gain output ratio was as high as 16, i.e. double the convention MSF-BR. The new NF-MSFDM configuration significantly reduced the unit's input thermal energy to suit the use of (the relatively expensive) solar energy as a desalination plant driver. On the other hand, the levelized water cost of NFMSF-DM (at TBT = 130 °C) is 14% lower than conventional MSF (at 110 ° C) at the current oil price 104 $/bbl [8]. In light of the recent development in the membrane filtration technologies, the cost of seawater pretreatment can be reduced if FO membranes were used instead of NF. The novel application of FO membrane for seawater filtration requires; firstly, retrofit FO system to the thermal desalination unit in a hybrid system. Secondly, to find a suitable draw solution that would maintain the low cost of FO pretreatment and reduces waste discharge to sea. Fortunately, the current FO membranes exhibit high water permeability and rejection rate which make them an ideal solution for seawater pretreatment. The objective of the current paper is to introduce a hybrid FO-thermal desalination system which is designed to remove scale elements from sweater to the thermal units and hence reduces scale deposition. The performance of the thermal evaporator will be evaluated after introducing the FO pretreatment. The scale deposition on the thermal unit will be estimated by using special software to predict the precipitation on inversely soluble metal ions on the heat exchangers. 2. Conceptual design of FO-thermal hybrid system The main purpose of FO-Thermal hybrid system is to alleviate the scale problem in the thermal desalination plants and hence reduces the system shutdown period. In the pretreatment process, FO membrane is

in charge of seawater softening i.e. the removal of divalent ions which are responsible of scale deposition on the heat exchangers. As discussed before, the driving force in FO is the osmotic pressure gradients between the draw and feed solutions. Different compounds such as ammonium carbon dioxide, magnesium sulfate, magnesium chloride, sodium chloride and many other chemicals were used as draw solutions. The cost of the draw solution will be added to the overall cost of FO treatment. Reverse ions diffusion and osmotic agent losses during the regeneration process will increase the cost of FO treatment [9,10]. The regeneration process can't guarantee a complete recycle of the osmotic agent and residue of ionic species, such as ammonium, in the product water is extremely undesirable [9]. Therefore, a careful selection of the draw solution will determine the success of FO seawater pretreatment from thermal desalination processes. Accordingly, it is proposed here to use the brine reject from the thermal desalination plants as a draw solution for the following reasons: 1. to reduce the cost and chemicals use in the FO process; 2. to reduce the amount of brine waste to discharge to sea; and 3. to reduce the intake seawater feed. The implementation of the new FO–MSF/MED design concept will increase the attractiveness of the FO pretreatment process. In the present paper, a two novel configurations of hybrid FO with MSF and MED respectively are developed and proposed as shown in Figs. 1 and 2. In case of FO–MSF hybrid system, it should be noted here that most of the MSF plants are operated in brine recycle mode. In such design, the brine reject is recycled back to the evaporated after mixing with the make-up from seawater feed. The make-up ratio depends on the MSF capacity, TBT, and recovery rate. For example, for seawater feed has 3.5 g/L TDS, 110 °C TBT, and recovery rate 31% the estimated make-up ration is about 10% of the total feed flow. These data were estimated by the VDS program for seawater desalination [2,8]. Typically, the concentration of the brine solution from the MSF plant is 30% higher than the seawater; this provides enough osmotic pressure gradients for FO operation. Usually, the temperature of brine reject is higher than the seawater. This will increase the osmotic pressure of the draw solution according to the following equation: Π ¼ nCRT

ð1Þ

Π is the osmotic pressure (bar), n number of ions, C is the molar concentration of solution, R is the gas constant, and T is the temperature (Kelvin). Fig. 1 shows a schematic diagram of the FO–MSF system. There are two options to feed seawater into the FO–MSF hybrid system. In the first option; seawater goes to FO first then to the heat rejection unit in MSF while the second option is to feed seawater into the heat rejection unit before it goes into the FO membrane system. The former option suits our design better because the feed temperature of seawater will be higher after leaving the heat rejection unit and this will increase its osmotic pressure according to Eq. (1). Therefore, Seawater is fed into one side of the FO membrane while the MSF brine reject is fed on the other side of the membrane in a concurrent flow mode. Due to the osmotic pressure difference across the FO, fresh water will cross the membrane from the seawater to the brine reject (draw solution) side. After leaving the MSF, seawater is sent to the heat rejection unit in the MSF plant before it is discharge to the sea. The diluted draw solution from the FO system will be the feed to the MSF. Inside the MSF, the feed solution is heated up to 110 °C and fresh water is extracted by evaporation while the concentrated brine is recycled to the FO membrane. Fig. 2 shows a novel hybrid FO–MED configuration. The brine of the last effect is recycled and diluted with FO membrane permeation. The diluted brine is then used as make up feed to the evaporator. The

A. Altaee et al. / Desalination 326 (2013) 19–29

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Fig. 1. FO–MSF hybrid system.

merit of the proposed configuration is excluding the divalent ions and reducing the intake of seawater feed.

The coefficient of membrane permeability, Aw, is calculated from the following equation:

3. FO membrane modeling

Aw ¼

A wealth of literature is available to estimate water and salt flow in FO membranes [11–13]. In this study, a simplistic model to predict the performance of FO membrane was adopted. Classical water and salt permeability equation are used in the model development based on the following hypotheses:

Jw is the membrane flux (l/m2h), ΔP is the applied pressure (bar), and ΔΠ is the osmotic pressure (bar). Salt permeability coefficient, B, is calculated from the membrane rejection rate, Rj, and Jw as shown in Eq. (3):

1. The brine reject from MSF is the draw solution in FO. Therefore the concentration of draw solution, CDS, equals the concentration brine reject, CBR. 2. The recovery rate of MSF/MED is 33% of the total feed (MSF is operated on brine recycle mode). 3. The flow rates of draw and feed solutions in the FO membrane are equal. In addition, it is assumed here that the rejection rate of FO membrane is similar to that used in the previous literature [14–16]. Typical rejection rate of FO membranes for divalent ions is more than 97%.



Jw ðΔP−ΔπÞ

ð1−RjÞ Jw Rj

ð2Þ

ð3Þ

Fig. 3 shows a schematic diagram of the FO membrane. The concentration of the draw solution is estimated from the concentration of feed water to the MSF, CfMSF, divided by the concentration factor 1/(1-Re) as the following: C DS ¼

C fMSF ð1−ReÞ

Fig. 2. FO–MED hybrid system.

ð4Þ

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draw solution flow rate into the FO, QDS-in, and the permeate flow rate, Qp. Substituting QDS-out in Eq. (10): C DS−out ¼

Q DS−in C Q DS−in þ Q p f

ð11Þ

Rearrange Eq. (11) to calculate the concentration of CDS-out: C DS−out ¼ Fig. 3. Schematic diagram of the FO process.

Cf : Qp 1þ Q DS−in

ð12Þ

In Eq. (12), the term (Qp/QDS-in) is equal to the FO recovery rate when QDS-in equals the flow of feed solution in FO, Qf. In this case Eq. (12) can be written as the following: In Eq. (4), Re is the recovery rate. In FO membrane, the TDS of permeate is the ratio of salt flux to water flux in the membrane and salt flux is function of salt permeability coefficient according to the following equations: J Cp ¼ s Jw

ð5Þ

  J s ¼ B C f −in −C p

ð6Þ

BC f −in Jw þ Cp

ð7Þ

It should be noted here that salt diffusion from the feed to the draw solution side of the membrane will affect the concentration of feed solution. Practically, salt diffusion from the feed to the draw solution side of the FO membrane will increase the TDS of draw solution. Therefore, the final TDS of the draw solution is estimated from the following equation: C DS−in ¼ C DS0−in þ C p

Qp Q DS−in

ð8Þ

where CDS-in is the final concentration of draw solution (mg/L), CDS0-in is the initial concentration of feed solution (mg/L), Qp is the permeate flow rate (m3/h), and QDS-in is the draw solution flow rate (m3/h). Eq. (8) is a useful formula to calculate the overall concentration of the draw solution and it can be used to estimate the concentration of the draw solution out, CDS-out. In RO membranes the concentrate concentration can be estimated as a function of the feed concentration, Cf, and the recovery rate, Re, according to the following equation: Cc ¼

Cf 1−Re

ð9Þ

In Eq. (9), Re is the ration of the permeate flow, Qp, to the concentrate flow, Qf. But Qf is the sum of Qp and the concentration flow, Qc. Therefore Eq. (9) can be rearranged to calculate the concentrate concentration: Cc ¼

Qf C Qc f

Cf : 1þR

ð13Þ

Eqs. (1)–(13) can be used to calculate the flow and concentrations of draw and feed solutions in the FO membrane. 4. Model development

where Js is salt flux (m/d), Cf-in is the concentration of seawater feed to the FO (mg/L), and Cp is the permeate concentration (mg/L). Substituting Eq. (6) in 5 and rearrange the equation to calculate the Cp: Cp ¼

C DS−out ¼

ð10Þ

Eq. (10) can be applied to calculate the concentration of draw solution out of the FO membrane, CDS-out. On contrary to RO membranes, the flow rate of draw solution out of the FO, QDS-out, is the sum of the

FO hybridization with thermal processes is a novel application to reduce the concentration of multivalent ions in the feed solution to the thermal desalination processes. Therefore, it is important to use a suitable membrane which exhibits a high water permeability and satisfactory rejection rate to multivalent ions such as Ca2+, Mg2+, and SO2− 4 . In the current process the rejection rate of FO membrane to NaCl is not an issue since it doesn't interfere with the mechanisms of scale deposition in the thermal processes but the high rejection to multivalent ions is important. Most importantly in FO pretreatment is the selection of membrane that it possesses high water permeability. Recovery rates between 16% and 32% were evaluated in the present study to simulate the situation of using FO membranes with different water permeabilities. The higher the membrane permeability the higher is its recovery rate. Table 1 shows the steps used to predict the concentration of diluted draw solution and estimated permeate flow for each recovery rate. 5. Scale deposition and model development The main ionic species which are responsible for scale deposition in MED and MSF are calcium carbonate, calcium sulfate and magnesium hydroxide. Scale deposition in thermal desalination processes

Table 1 Procedure to estimate the FO membrane performance.

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Fig. 4. Chemical relations of scale deposition in thermal desalination.

is determined by the TBT of the desalination plant. During the process of scale deposition, different chemical reactions take place in the thermal plant as shown in Fig. 4. In this study, a model for calcium carbonate scale formation is developed. Calcium carbonate scale or alkaline scale formation was reported in MED plants due to the CO2 release from feed solution. The TBT in MED is less than 70 °C, each MED contains a number of consecutive effects which work at different temperatures and pressures. Inside the each effect there is a tube bundle which is divided into elements (rows) (Fig. 5). Two models were pursued to predict scale deposition on MED elements; i.e. CaCO3 scale precipitation model and CO2 release model.

Obviously, from the equation above, CaCO3 formation and deposition is a function of CO2 release. It is widely accepted now that the growth of scale layer is a net result of deposition and removal processes [17]. Whilst scale deposition is a continuous process, the removal process is negligible especially when the scale layer consists of a pure crystalline salt. The thermal fouling rate is therefore determined by the deposition flux, ω, of the crystallizing layer:

5.1. CaCO3 scale precipitation model

In Eq. (15), Rf is the thermal resistance of scale layer (m2K/W), kf is the thermal conductivity of the scale layer (W/m.°C), ρf is the scale density (kg/m3). It was found that the crystallization flux of CaCO3 on the scale-water interface is governed by the following kinetic expression [17,18]:

The deposition of a CaCO3 layer on the heated tubes can simply be described by the overall crystallization reaction as in the following equation: 2þ

Ca



þ 2HCO3 →CaCO3 ðsÞ þ CO2 ðg Þ þ H 2 O

ð14Þ

dR f 1 dx ω ¼ ¼ k f dt ρ f k f dt

ω ¼ kR

ð15Þ

nh ih i o 2þ 2− Ca i CO3 −K sp

ð16Þ

where the subscript i refers to the actual concentrations on the reaction surface. The crystallization rate coefficient, kR (m4/kg.s), is an empirical value and it is specific to a crystallizing salt which is affected by the presence of substances interfering with crystal growth. It is important to mention here that kR is a temperature dependent factor according to the Arrhenius relationship [17,18]. In Eq. (16), Ksp is the solubility product constant of calcite in seawater and it is given by the following equation [18]:   2903:293 þ 71:595 logðt Þ K sp ¼ anti: log −171:9773−0:077993T þ T ð17Þ

Table 2 Ryznar Scale Index.

Fig. 5. The liquid distribution in the evaporator.

RSI N 9 7.5 b RSI b 9 6 b RSI b 7 5 b RSI b 6 4 b RSI b 5

Very intensive dissolving of scale and corrosion Intensive dissolving of scale and corrosion Stable water, slight tendency for dissolving of scale moderate to slight scaling Severe Scaling

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Fig. 6. Schematic diagram of scale model.

Desorption rate for CO2 release, in absence of chemical reaction, is given by the following equation:

5.2. CO2 release model There are a number of assumptions to be made for the calculation of the CO2 release: - The crystallization and CO2 diffusion is considered as the limiting phenomenon. - The liquid film is supposed to be thin. - The evaporator is under vacuum conditions.

0

NCO2 ¼ K L ½CO2 m

ð18Þ

where [CO2] m is the average concentration of the CO2 in the element, NCO2 is the deposition rate of CO2, and K0L is the mass transfer coefficient.

Fig. 7. Interface of VDS software of FO–MSF.

A. Altaee et al. / Desalination 326 (2013) 19–29

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Fig. 8. Interface of VDS software of FO–MED.

Finally, Ryznar Scale Index (RSI) can be calculated from the following equation: RSI ¼ 2:pHs−pH

ð19Þ

where pHs is the pH at saturation in calcite or calcium carbonate, and pH is the measured water pH. RSI less than 5 indicates a severe scale problem while 6 b RSI b 7 indicates a stable water with slight tendency for dissolving scale (Table 2). 5.3. Description of the model Fig. 6 shows the steps followed in this study for scale modeling in MED. In the first step all input parameters are fixed and the model calculates seawater composition at the entrance of the evaporator (first element) in step two. CO2 release in the element and the average scale thickness are calculated in step three. Furthermore, the model calculates the quantity of the vapor produced in the element. Considering all these phenomena (CO2 release, scale deposition and evaporation), the model estimates the properties of seawater at the exit of the element i which corresponds to the entrance of the element i + 1.

VDS package is developed for the design and simulation of different types and configurations of the thermal and membrane desalination processes. Object oriented programming with Visual Basic (VB) is used to offer a flexible, reliable, and friendly user-interface. The interface, as shown in Figs. 7 and 8, aids the designer to perform different calculations such as mass and heat balance, exergy, and thermo-economics. In addition, the package enables the designer to perform different modifications to an existing plant or to develop the conceptual design of new configurations. The VDS is utilized to perform process design calculations specifying the heating steam operating conditions (pressure, temperature), the target capacity of the evaporator (distillate rate per hour), TBT, sea water conditions (temperature, and salinity), make-up flow rate, brine recirculation salinity, blow down and reject brine temperature. Some design parameters such as number of stages, tube length, diameters, material type are also calculated. In the present study, both configuration FO–MSF and FO–MED are simulated under different recovery ratio of FO membrane (16 % up to 32%). Some of the operating parameters such as (heating steam

Table 4 Parameters value in FO-thermal hybrid system.

6. Thermal process simulation In this study, the developed Visual Design and Simulation (VDS) software, [19–22], is utilized as a reliable techno-economical tool to predict the performance of FO-thermal processes. The present

Parameter

Value

Brine Reject TDS Seawater TDS Brine Reject temperature Seawater temperature Brine Reject Osmotic Pressure⁎ Seawater Osmotic Pressure⁎ Thermal plant Recovery Rate FO Recovery Rate

67,203 ppm 45,000 ppm 44 °C 30 °C 52.5 bar 33.6 bar 33% 16%–32%

⁎ Van't Hoff Equation (Π = nCRT).

Table 3 Seawater composition. Ions

Concentration (ppm)

K Na Mg Ca Cl SO4 HCO3 TDS

496 13,812 1657 539 24,868 3472 182 45,026

Table 5 Water and salt permeability in FO membrane. %Re

Jw (L/m2h)

Cp (mg/L)

Js kg/m2h

16 20 24 28 32

9.6 8.3 6.9 5.5 4.0

891 914 938 965 995

0.0086 0.0076 0.0065 0.0053 0.0040

26

A. Altaee et al. / Desalination 326 (2013) 19–29

Fig. 9. Ca concentrations in MSF stages 1 to 19.

conditions, TBT, recycle flow rate, recycle TDS, number of stages) are kept the same. The distillate flow rate, process streams flow rate and temperature are calculated using VDS. 7. Analysis of the results 7.1. FO–MSF Plant The salinity of seawater used in this study was taken from the previous literature [9]. Seawater salinity may vary from 35,000 to 48,000 ppm depending on the location. In this study a salinity of 45,000 ppm was considered to resemble seawater salinity in the Middle East where thermal desalination processes are popular. The composition of seawater is illustrated in Table 3, which shows the

concentrations of major ions in seawater. Seawater and brine reject from the thermal plant are, respectively, the feed and draw solution for the FO membrane system. Normal seawater temperature in the Gulf Region of the Middle East varies from 15 °C to 35 °C. The temperature of seawater was fixed at 35 °C in this study. Table 4 shows the simulation parameters for FO–MSF hybrid system. It is worth to be mentioned here that Van't Hoff equation was used to calculate the osmotic pressure of seawater and the brine solution. The total number of stages in the MSF plant was 21 stage; 19 of which are in the heat recovery unit and 2 stages in the heat rejection unit. To reduce the cost of FO treatment, brine reject from the thermal plant is used as a draw solution in the FO membrane. Due to the osmotic gradients, fresh water will cross from the seawater to the draw solution

Fig. 10. Mg concentrations in MSF stages 1–19.

A. Altaee et al. / Desalination 326 (2013) 19–29

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Fig. 11. SO4 concentration in MSF stages 1 to 19.

side of the FO membrane and dilutes the draw solution. The diluted draw solution is then sent to the MSF plant as a feed solution. Eq. (12) was used to calculate the TDS of draw solution and it was 58,056, 56,155, 54,378, 52,714, and 51,153 ppm respectively for the FO recovery rates 16%, 20%, 24%, 28% and 32%. The water and salt permeability in FO membrane is shown in Table 5. Simulation results showed that the TDS of the feed solution to the MSF plant decreased with increasing the recovery rate of FO membrane. Scale problem in MSF plant is mainly due to the precipitation of MgSO4 and CaSO4 in the heat exchanger. This type of non-alkaline scale is more severe in the brine heater of the MSF due to the high operating temperature (TBT). Figs. 9–11 show the concentrations of Ca, Mg, and SO4 in the flashing brine through MSF stages 1 to 19. The higher the FO recovery rate the lower the concentrations of Ca, Mg and SO4 are in the feed solution. Primarily, this was due to the higher dilution of the draw solution at high FO recovery rate. For

example, the concentration of Ca in feed was 698 and 614 ppm for 16% and 32% FO recovery rates. For the same recovery rates, the concentration of Mg was 2146 and 1888 ppm while the concentration of SO4 was 4398 and 3956 ppm respectively. A low concentration of Ca, Mg and SO4 in the brine heater and the high temperature stages of the MSF is extremely desirable due to the high tendency of scale development in these stages. Compared to seawater, up to 23% reduction in the concentration of Ca, Mg and SO4 was achieved at 32% FO recovery rate. It was also observed from the simulation results that the MSF distillate flow rate increased with increasing the FO recovery rate. This was due to the higher feed solution to MSF at higher FO recovery rates. Fig. 12 shows the distilled water recovery rate at different FO recovery rates. As indicated, the higher FO recovery rate the higher was the distilled water generated in the thermal evaporators.

Fig. 12. Distillate produced in MSF stages 1 to 19 at different FO recovery rates.

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A. Altaee et al. / Desalination 326 (2013) 19–29 Table 6 RSI of feed solution for different FO recovery rates. Re(%)

RSI

0 16 20 24 28 32

4.67 4.80 4.79 4.82 4.85 4.88

Table 7 The TDS of brine reject before and after dilution. Cbefore ppm

K Na Mg Ca Cl SO4 HCO3 CO3 ΣTDS

708.6 19,731.4 2367.1 770.0 35,525.7 4960.0 260.0 12.7 64,336

16% Re

20%Re

24%Re

28% Re

32% Re

Cafter ppm

Cafter ppm

Cafter ppm

Cafter ppm

Cafter ppm

306.9 19,571.3 1021.6 332.3 32,380.3 2140.5 112.6 15.5 55,881

297.1 18,930.3 987.9 321.3 31,321.3 2069.9 109.0 13.7 54,050

287.9 18,331.2 956.3 311.1 30,331.9 2003.9 105.6 13.2 52,341

279.4 17,770.3 926.8 301.5 29,405.5 1942.0 102.5 12.7 50,728

271.3 17,244.1 899.1 292.5 28,536.6 1883.9 99.6 12.2 49,239

7.2. FO–MED plant The typical operation mechanism of the MED process is slightly different to the MSF in which the brine concentrate is entirely discharged to seawater without recycling. With a recovery rate slightly over 30%, the brine concentrate from the MED needs more than 30% dilution to reach the normal ionic concentrations of seawater. RSI of feed solution for different FO recovery rates was calculated from Eq. (19) and listed in Table 6. For all FO recovery rates the RSI of feed solution was less than 5 which indicates a sever scale problems (Table 2). Therefore, the recovery rate of FO should be high enough to reduce the concentration of feed water to MED below the concentrations of seawater and hence justify the FO pretreatment. Alternatively, it is proposed here to dilute the brine reject from MED with NaCl stock solution before entering the FO membrane to reduce the concentration of the

900

5.6

800

5.4

700 5.2 600 500

4.8

400

RSI

5

300 4.6 200

RSI

4.4

100

Ca concentration (mg/L) 4.2

0 0

16

20

24

28

%Re Fig. 13. RSI and Ca concentrations at different FO recovery rates.

32

Ca concentration (mg/L)

Ions

divalent ions. For 45,000 mg/L seawater salinity, VDS program estimated the TDS of brine reject from the MED is around 65,000 mg/L (Fig. 8). After leaving the MED, the brine reject is diluted (only once) with NaCl solution of equal TDS concentration to reduce the concentration of divalent ions. Then it goes to the FO membrane system for further treatment. Table 7 shows the concentrations of brine reject before and after FO treatment. As shown in 7, the concentration of Ca ions, which is the main element in CaCO3 scale layer, has dropped to half after dilution with NaCl but it didn't change significantly with increasing the FO recovery rates from 16% to 32%. At 16% and 32% FO recovery rates the concentrations of Ca ion was 332 ppm and 292 ppm respectively. Ryznar Stability Index (RSI) was calculated for the different feed solutions to MED with and without FO pretreatment. Using the scale model described in Section 5, a sharp drop in the scale tendency was observed when feed solution to MED was pretreated by FO membranes (Fig. 13). Without FO pretreatment, the RSI of the feed water to MED was 4.6 then increased to 5.5 when 16% recovery rate FO pretreatment was applied. It is worth mentioning here that 4 b RSI b 5 indicates a sever scale tendency while 5 b RSI b 6 indicates a moderate to slight scale tendency. As such the FO pretreatment can successfully reduce the scale tendency of the MED process. However, there wasn't a significant difference in the MED scale tendency when FO pretreatment increased from 16% to 32% due to the slight decrease in the concentration of Ca in the MED feed solution (Fig. 13). In the first effect of the MED plant, the thickness of CaCO3 scale layer on the bottom tube rows was 43 μm and 39 μm for 20% and 32% FO recovery rates respectively 9 Fig. 11). Without FO pretreatment, however, the thickness of the CaCO3 scale layer was increased to 74 μm. This indicates to the effectiveness of FO pretreatment in the removal of Ca ions from feed solution to MED and hence reducing the thickness of CaCO3 scale layer. Fig. 11 shows that the thickness of CaCO3 was the highest in effect number 1 and decreased towards effect number 6. Primarily, this was due to the high operating temperature in effect number 1 which gradually decreased towards the last effect. Fig. 14 also shows that the thickness of CaCO3 scale layer on the top tube rows. The profile of CaCO3 scale deposition was found to be similar to that in the bottom tube rows. The highest CaCO3 thickness was in the first effect then decreased towards the last effect due to the gradual drop in the operating temperature from effect 1 to 6. Furthermore, the thickness of CaCO3 in the top tube rows was much lower than in the bottom tube rows. These results are in agreement with the previous studies which also found that the thickness of CaCo3 was the highest

A. Altaee et al. / Desalination 326 (2013) 19–29 Top row without FO treatment Top row FO Re=20% Top row FO Re=32% Bottom row without FO treatment Bottom row FO Re=20% Bottom row FO Re=32%

80 70

CaCO3 thickness (µm)

29

60 50 40 30 20 10 0 0

1

2

3

4

5

6

7

MED effect No. Fig. 14. Thickness of the scale layer on the tube bundles in MED.

in the bottom tube rows [18,21]. In the present study, FO pretreatment demonstrated a high potential of reducing the scale tendency in the thermal desalination processes; this in turn will reduce shut-down time and chemicals use in the thermal processes. 8. Conclusion - Two novel hybrid FO–MSF and FO–MED configurations are proposed and simulated in this study. The main role of FO membrane was to remove multivalent ions from the feed water to the thermal desalination plants. - The simulation results showed that the FO pretreatment successfully reduced the concentration of multivalent ions which are responsible for the scale deposition in MSF and MED. The results also showed that the removal of divalent ions increased with increasing the recovery rates of FO. It is, therefore, desirable to use high permeability membranes in the FO pretreatment of seawater. - A scale model was developed to estimate the thickness of CaCO3 layer on the tube bundles in the MED plant. The results showed a higher scale tendency in the bottom tube rows than in the top tube rows. The results also showed that the thickness of CaCO3 layer decreased in the direction from effect 1 to 6 due to the lower operating temperatures. - The FO pretreatment for feed solution to the thermal desalination (MSF/MED) processes will also help to increase the TBT and hence the recovery rate which is recommended to be investigated in the future work by the authors. - The proposed configurations helps to reduce brine reject discharge to sea and chemicals use and to reduce the seawater feed. - The simulation results revealed that a high water permeability and divalent rejection rate membrane is preferably used in the FO pretreatment. Further research should be carried out to investigate this issue. References [1] Abdel Nasser Mabrouk, Techno-economic Analysis of Seawater Desalination Plants, LAP LAMBERT Academic Publishing, Germany, 2012, ISBN 978-3-659-17934-1. [2] A. Mabrouk, Techno-economic analysis of once through long tube MSF process for high capacity desalination plants, Desalination 315 (2013) 84–94. [3] Atta. Hassan, Fully integrated NF-thermal seawater desalination process and equipment, US Patents No 2006/0157410 A1, July 20, 2006.

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