**Techno-Economics of Hybrid NF/FO with Thermal Desalination Plants**

Abdel Nasser Mabrouk, Hassan Fath, Mohamed Darwish and Hassan Abdulrahim

Additional information is available at the end of the chapter

http://dx.doi.org/10.5772/60207

#### **Abstract**

Hybrid desalination technology is defined as any combination of thermal and membrane processes in seawater desalination systems. So far, the two technologies have evolved rather independently with some degree of competition. Traditionally, in co-generation market applications, thermal desalination has succeeded in estab‐ lishing a stronghold where large capacities are needed, energy costs are low, and seawater quality is challenging. However, in recent years, membrane systems have also succeeded in grabbing a larger share of the world seawater desalination market, mainly as a result of progress made in membrane and energy recovery technologies. Realizing the potential benefits and challenges faced by both technologies on their own, designers have been looking for ways to synergize and combine the two technologies in optimum configurations, which promise to further reduce the total cost of seawater desalination.

Several studies have been published over the past 20 years addressing the potential of integrating hybrid desalination systems. Coupling schemes worth noting for hybrid systems include RO preheating using condenser heat reject of the associated distilla‐ tion unit; the use of membrane filtration upstream of MSF, MED, and RO systems; brine recirculation coupling; product blending; and the use of common intake and outfall systems. To date, commercially available hybrid desalination plants are of the simple non-integrated type. They may share common systems such as intake and outfall facilities, but otherwise they run independently at the same site. Product water from the membrane and thermal systems are usually blended to international standards on water quality. One more step ahead this chapter addresses the role of

© 2015 The Author(s). Licensee InTech. This chapter is distributed under the terms of the Creative Commons Attribution License (http://creativecommons.org/licenses/by/3.0), which permits unrestricted use, distribution, and reproduction in any medium, provided the original work is properly cited.

using forward osmosis (FO) or nano-filtration (NF) as a pre-treated method to the existing thermal desalination plants. The target of this hybridization is to reduce divalent ions that cause hard scale deposition at elevated temperature. The separation of divalent ion enables the increase of the desalination process temperature greater than 110°C, which consequently increases the plant performance, increases the productivity, and reduces the chemical consumption.

Integrating the NF system with new (MSF-DBM) configuration enables to increase the TBT up to 130 °C. The new NF-MSF-DBM configuration significantly reduces the unit's input thermal energy to suit the use of (the relatively expensive) solar energy as a desalination plant driver. The desalination pilot test is built to evaluate the performance of the novel de-aeration and brine mix (MSF-DBM) configuration at high top brine temperature (TBT) using NF membrane. The capacity of the desalination pilot plant is 1.0 m3 /day of water. Comparisons between the simulation and the experimental results of the pilot unit subsystems are relatively satisfactory. The newly developed NF-MSF-DBM (de-aerator and brine mix) configuration is tested at TBT = 100 °C. The gain output ratio (GOR) is calculated as 15, which is almost twice the GOR of the traditional MSF. High GOR NF-MSF-DBM system requires lower input thermal energy, which makes the integration with (the relatively expensive) RE system as a desalination plant driver, viable option.

Simulation results showed that integrating FO to existing MSF and using brine of the last stage as a draw solution at a recovery ratio of 35% reduce the Ca+ ions in the seawater feed by 20%, which enables increasing the TBT up to 130°C safely. The simulation results show that, the production of the existing MSF plant increases by 20% as a result of working at higher temperature of TBT=130 °C. The specific operating cost (OPEX) analysis showed that up to 2.3 M\$/year of chemical cost can be saved if the FO membrane is deployed to the existing MSF desalination plants in Qatar.

**Keywords:** Desalination, Thermal, Membrane, Hybrid, Cost

### **1. Introduction**

#### **1.1. Hybrid desalination concept**

Hybrid technology is defined as any combination of thermal and membrane processes in seawater desalination systems. So far, the two technologies have evolved rather independ‐ ently with some degree of competition. Traditionally, in co-generation market applica‐ tions, thermal desalination has succeeded in establishing a stronghold where large capacities are needed, energy costs are low, and seawater quality is challenging. However, in recent years, membrane systems have also succeeded in grabbing a larger share of the world seawater desalination market, mainly as a result of the progress made in membrane and energy recovery technologies. Realizing the potential benefits and challenges faced by both

technologies on their own, designers have been looking for ways to synergize and combine the two technologies in optimum configurations, which promise to further reduce the total cost of seawater desalination.

Several studies have been published over the past 20 [1-14] years addressing the potential of integrating hybrid desalination systems. Coupling schemes worth noting for hybrid systems include RO preheating using condenser heat reject of the associated distillation unit; adding thermal vapour compression to MED systems; the use of membrane filtration (NF) upstream of MSF, MED, and RO systems; brine recirculation coupling; product blending; and the use of common intake and outfall systems.

To date, commercially available hybrid plants are of the simple non-integrated type. They may share common systems such as intake and outfall facilities, but otherwise they run independ‐ ently at the same site. Product water of both membrane and thermal plants are usually blended to meet the international standards water quality specifications. Examples of existing hybrid plants include Jeddah, Al-Jubail, and Yanbu in Saudi Arabia. Recently, a power and water plant has been designed and built by Doosan in Fujairah (UAE). This plant produces 500 MW of net electricity to the grid and 100,000 MIGD of fresh water, 63% of which is produced by MSF with the balance produced by RO [1].

### **1.2. NF's role in desalination**

using forward osmosis (FO) or nano-filtration (NF) as a pre-treated method to the existing thermal desalination plants. The target of this hybridization is to reduce divalent ions that cause hard scale deposition at elevated temperature. The separation of divalent ion enables the increase of the desalination process temperature greater than 110°C, which consequently increases the plant performance, increases the

Integrating the NF system with new (MSF-DBM) configuration enables to increase the TBT up to 130 °C. The new NF-MSF-DBM configuration significantly reduces the unit's input thermal energy to suit the use of (the relatively expensive) solar energy as a desalination plant driver. The desalination pilot test is built to evaluate the performance of the novel de-aeration and brine mix (MSF-DBM) configuration at high top brine temperature (TBT) using NF membrane. The capacity of the desalination

experimental results of the pilot unit subsystems are relatively satisfactory. The newly developed NF-MSF-DBM (de-aerator and brine mix) configuration is tested at TBT = 100 °C. The gain output ratio (GOR) is calculated as 15, which is almost twice the GOR of the traditional MSF. High GOR NF-MSF-DBM system requires lower input thermal energy, which makes the integration with (the relatively expensive) RE system as a

Simulation results showed that integrating FO to existing MSF and using brine of the last stage as a draw solution at a recovery ratio of 35% reduce the Ca+ ions in the seawater feed by 20%, which enables increasing the TBT up to 130°C safely. The simulation results show that, the production of the existing MSF plant increases by 20% as a result of working at higher temperature of TBT=130 °C. The specific operating cost (OPEX) analysis showed that up to 2.3 M\$/year of chemical cost can be saved if the FO membrane is deployed to the existing MSF desalination plants in Qatar.

Hybrid technology is defined as any combination of thermal and membrane processes in seawater desalination systems. So far, the two technologies have evolved rather independ‐ ently with some degree of competition. Traditionally, in co-generation market applica‐ tions, thermal desalination has succeeded in establishing a stronghold where large capacities are needed, energy costs are low, and seawater quality is challenging. However, in recent years, membrane systems have also succeeded in grabbing a larger share of the world seawater desalination market, mainly as a result of the progress made in membrane and energy recovery technologies. Realizing the potential benefits and challenges faced by both

/day of water. Comparisons between the simulation and the

productivity, and reduces the chemical consumption.

pilot plant is 1.0 m3

216 Desalination Updates

**1. Introduction**

**1.1. Hybrid desalination concept**

desalination plant driver, viable option.

**Keywords:** Desalination, Thermal, Membrane, Hybrid, Cost

NF, in particular, has been advocated as a pretreatment option upstream of a thermal desali‐ nation unit [2-8]. Due to small pore size and charge at the surface of the membrane, NF is known to remove divalent ions, including a fraction of scale-producing hardness and salts, allowing in principle at least a possible increase in top brine temperature and promising improved steam economy. Studies on NF-MSF pilot tests claim scale-free operation for 1,200 hours with top brine temperatures reaching 130°C, reporting an improvement in recovery from 30% for stand-alone MSF to 70% with NF [2-8]. The integration of NF –MED system is under pilot investigation by Saline Water Desalination Research Institute (SWDRI) with Sasakura (a Japanese-based consortium) [15]. Implementation of dual-stage NF has been successful evaluated at Long Beach, California [16]. From the present analysis, NF will play a crucial role in desalination, provided that the cost of NF membrane would be decreased. Efficient removal of boron has also been reported [16]. However, the reliability and economic viability of such a design need to be confirmed, considering the higher thermal and pressure load implied by the design and the additional capital, energy, and operation and maintenance costs of membrane pre-treatment components. Adding to the uncertainty are increased risks for corrosion and the long-term reliability of such a system.

#### **1.3. FO role in desalination**

The role of using forward osmosis (FO) as a pre-treated method to the existing thermal desalination MSF/MED plants is to reduce divalent ions that cause hard-scale deposition at elevated temperature. The removal of the divalent ions, such as CaSO4, from the MSF feed enables to increase the desalination process temperature greater than 110°C. Consequently the plant performance and productivity will increase. Due to the removal of the ions which cause scale deposit, the chemical additive consumption will be decreases. In the MSF process due to working at higher temperature, hard scale, such as calcium sulphate, is formed. As calcium sulphate is two orders of magnitude more soluble than calcium carbonate, the sulphate is much less likely to drop out of solution when both are present.

In the light of the recent development in the membrane filtration technologies, the cost of seawater pretreatment can be reduced if FO membranes were used with/without NF. The novel application of FO membrane for seawater filtration requires, firstly, retrofitting the FO system to the thermal desalination unit. Secondly, it also requires finding a suitable draw solution that would reduce the cost of FO pre-treatment. Fortunately, the current FO membranes exhibit high water permeability and rejection rate, which make them an ideal solution for seawater pretreatment [17]. A novel hybrid FO-thermal desalination system to remove scale deposit elements from seawater to the thermal units, is presented [17-18]. The performance of the thermal evaporator was evaluated after introducing the FO pretreatment. The scale deposition on the thermal unit was estimated by using special software to predict the precipitation on inversely soluble metal ions on the heat exchangers [17-18].

#### **1.4. Objective**

This chapter addresses the role of using FO or NF as a pre-treated method to the existing thermal desalination plants. The target of this hybridization is to reduce divalent ions that cause hard-scale deposition at elevated temperature. The separation of divalent ion enables the increase of the desalination process temperature greater than 110°C, which consequently increases plant performance and productivity, as well as reduces the chemical consumption.

### **2. Process description**

#### **2.1. MSF-RO hybrid**

The simple type refers to co-located thermal and membrane systems that may share some common systems on site. This in turn facilitate blended to product water specifications, but otherwise are running independently. Examples include the Fujairah plant and three Saudi plants in Jeddah, Al-Jubail, and Yanbu. The Fujairah plant [1], representing a simple hybrid type, was constructed by Doosan Heavy Industries and is currently considered the largest existing hybrid type. The plant is rated at 100 MIGD, of which 63% are produced by MSF and 37% by RO. Featuring a combined cycle system, it also generates 500 MW to the grid. The thermal part of the plant includes five MSF evaporators rated at 12.5 MIGD each, with a top brine temperature of 110◦C. The membrane part includes two RO passes, using a conventional pretreatment system and energy recovery devices of the Pelton type. A specification of 200 PPM as the maximum water product salinity was met by the design.

Another option of a hybrid type is to improve the membrane performance. This type includes the integration of hybrid membrane and thermal systems, with the aim of improving mem‐ brane recovery by preheating the RO feed using heat reject from the thermal unit as shown in Figure (1). Higher temperatures are known to improve membrane flux, mainly as a result of reduced viscosity. Several investigators examined the effects of preheating in pilot tests, and about 3% of recovery improvement is reported per degree Celsius [19]. This should, however, be weighed against potential negative effects of high temperatures on membrane performance, in particular compaction damage. Membrane manufacturers have traditionally set an upper temperature limit of 40◦C for the use of their membranes, and it is not clear how close to this limit operation should be, in order to optimize life-cycle membrane performance and costs. The measure is particularly useful in winter when seawater temperatures are reduced.

**Figure 1.** Process flow diagram of hybrid RO-MSF process

Figure 1: Process flow diagram of hybrid RO-MSF process Preheating the Fujairah RO feed in winter is an example of integrated hybrid operation, representing type 2. A 10°C increase from 23°C to 33°C for an RO unit equivalent in capacity to the Fujairah plant would increase recovery by about 30% and, therefore, reduce feed pressure requirements. This preheating feature could then be made to good use in the winter when seawater temperatures in the Gulf drop by 15–20°C.

**Cooling reject**

#### **2.2. NF-MSF process**

Figures

plant performance and productivity will increase. Due to the removal of the ions which cause scale deposit, the chemical additive consumption will be decreases. In the MSF process due to working at higher temperature, hard scale, such as calcium sulphate, is formed. As calcium sulphate is two orders of magnitude more soluble than calcium carbonate, the sulphate is much

In the light of the recent development in the membrane filtration technologies, the cost of seawater pretreatment can be reduced if FO membranes were used with/without NF. The novel application of FO membrane for seawater filtration requires, firstly, retrofitting the FO system to the thermal desalination unit. Secondly, it also requires finding a suitable draw solution that would reduce the cost of FO pre-treatment. Fortunately, the current FO membranes exhibit high water permeability and rejection rate, which make them an ideal solution for seawater pretreatment [17]. A novel hybrid FO-thermal desalination system to remove scale deposit elements from seawater to the thermal units, is presented [17-18]. The performance of the thermal evaporator was evaluated after introducing the FO pretreatment. The scale deposition on the thermal unit was estimated by using special software to predict the precipitation on

This chapter addresses the role of using FO or NF as a pre-treated method to the existing thermal desalination plants. The target of this hybridization is to reduce divalent ions that cause hard-scale deposition at elevated temperature. The separation of divalent ion enables the increase of the desalination process temperature greater than 110°C, which consequently increases plant performance and productivity, as well as reduces the chemical consumption.

The simple type refers to co-located thermal and membrane systems that may share some common systems on site. This in turn facilitate blended to product water specifications, but otherwise are running independently. Examples include the Fujairah plant and three Saudi plants in Jeddah, Al-Jubail, and Yanbu. The Fujairah plant [1], representing a simple hybrid type, was constructed by Doosan Heavy Industries and is currently considered the largest existing hybrid type. The plant is rated at 100 MIGD, of which 63% are produced by MSF and 37% by RO. Featuring a combined cycle system, it also generates 500 MW to the grid. The thermal part of the plant includes five MSF evaporators rated at 12.5 MIGD each, with a top brine temperature of 110◦C. The membrane part includes two RO passes, using a conventional pretreatment system and energy recovery devices of the Pelton type. A specification of 200

Another option of a hybrid type is to improve the membrane performance. This type includes the integration of hybrid membrane and thermal systems, with the aim of improving mem‐

PPM as the maximum water product salinity was met by the design.

less likely to drop out of solution when both are present.

inversely soluble metal ions on the heat exchangers [17-18].

**1.4. Objective**

218 Desalination Updates

**2. Process description**

**2.1. MSF-RO hybrid**

B/H **Rejection NF membrane Seawater Feed Blow down NF Permeate Brine blow down Permeate Brine recycle MSF recovery section Distillate** This type includes the integration of hybrid membrane and thermal systems with the objective of improving the gain output ratio (GOR) and steam economy of the thermal system (see Figure 2). The GOR is a function of the available temperature range and can, therefore, be improved by increasing the top brine temperature (TBT). Traditionally, the top brine temperature is limited to 110 °C for MSF and is limited to 65◦C for MED. This practice, in addition to chemical dosing and mechanical cleaning, is necessary to minimize scale deposition on heat transfer surfaces. Some investigators have advocated the use of NF membrane upstream of the thermal system as a pretreatment step to reduce scaling hardness and additionally some salt. This would, in principle, allow operation at higher temperatures, without increased scaling. SWCC investigators have tested a hybrid NF/MSF pilot unit running at a top brine temperature of 130°C for a period of 1,200 hours and reported a doubling in the recovery with no observed scale formation [10].

Figure 2: Process flow diagram of hybrid NF-MSF process.

Figures

Figure 2: Process flow diagram of hybrid NF-MSF process.

**Brine blow down**

**Cooling reject**

**Brine recycle**

Figure 1: Process flow diagram of hybrid RO-MSF process

**RO membrane**

**RO Permeate**

**Distillate Blend prodcut**

Desalinated water

Cooling reject

Seawater

*W, permeate*

*S, permeate*

**Permeate**

**MSF**

**Seawater Feed Blow down**

**Figure 2.** Process flow diagram of hybrid NF-MSF process

E-3 **Rejection**

**MSF recovery section**

#### **2.3. FO-MSF**

Figure 3 shows hybrid FO-multi stage flash (FO-MSF) system for high TBT MSF. In this type of hybrid system, the brine reject from the thermal desalination process will be considered as a draw solution, while the cooling seawater exiting from the MSF heat rejection section will be used as a feed solution. Permeate water will transport across the FO membrane from the feed to the draw solution side while monovalent and multivalent ions are rejected by the FO membrane. After leaving the FO membrane, the concentrated seawater is dumped back to the sea. Simultaneously, the diluted draw solution from the FO process is circulated to the MSF recovery section. Inside the MSF plant, fresh water is extracted from the draw solution by evaporation and is condensed in the consecutive MSF chambers. The distilled water is collected at the last stage and directed to the distilled tank. The un-flashed brine through MSF stages (brine pool) is collected in the last stage of MSF evaporator at high salinity and then is directed to the FO as a draw solution.

FO membrane

Figure 4: Process flow diagram of hybrid FO-MED process

*W, brine*

*S, brine*

Figure 5: Schematic diagram of the NF membrane streams.

MED Feed

Condneser

Vapor

Figure 3: Process flow diagram of hybrid FO-MSF process **Figure 3.** Process flow diagram of hybrid FO-MSF process

MED

Last cell

Condensate return

*W, feed S, feed*

Steam

Desalinated water

Cooling reject

**Cooling reject**

#### **2.4. FO-MED** B/H **Rejection FO membrane**

Figures

E-3 **Rejection**

B/H **Rejection**

**Distillate**

**MSF recovery section**

**Figure 2.** Process flow diagram of hybrid NF-MSF process

**2.3. FO-MSF**

220 Desalination Updates

to the FO as a draw solution.

**MSF recovery section**

**Figure 3.** Process flow diagram of hybrid FO-MSF process

MED

Last cell

Condensate return

*W, feed S, feed*

Steam

**MSF recovery section**

**RO membrane**

**NF membrane**

**NF Permeate**

**Seawater Feed**

Desalinated water

Cooling reject

Seawater

*W, permeate*

*S, permeate*

**Permeate**

**Cooling reject**

**Diluted brine**

**Permeate**

**Brine blow down**

**Seawater Feed Blow down**

**RO Permeate**

**Distillate Blend prodcut**

**Permeate**

**MSF**

**Brine blow down**

**Cooling reject**

**Brine recycle**

Figure 1: Process flow diagram of hybrid RO-MSF process

**Cooling reject**

**Brine recycle**

Figure 2: Process flow diagram of hybrid NF-MSF process.

Figure 3 shows hybrid FO-multi stage flash (FO-MSF) system for high TBT MSF. In this type of hybrid system, the brine reject from the thermal desalination process will be considered as a draw solution, while the cooling seawater exiting from the MSF heat rejection section will be used as a feed solution. Permeate water will transport across the FO membrane from the feed to the draw solution side while monovalent and multivalent ions are rejected by the FO membrane. After leaving the FO membrane, the concentrated seawater is dumped back to the sea. Simultaneously, the diluted draw solution from the FO process is circulated to the MSF recovery section. Inside the MSF plant, fresh water is extracted from the draw solution by evaporation and is condensed in the consecutive MSF chambers. The distilled water is collected at the last stage and directed to the distilled tank. The un-flashed brine through MSF stages (brine pool) is collected in the last stage of MSF evaporator at high salinity and then is directed

B/H **Rejection FO membrane**

**Diluted brine**

Figure 3: Process flow diagram of hybrid FO-MSF process

MED Feed

Condneser

Vapor

FO membrane

Figure 4: Process flow diagram of hybrid FO-MED process

*W, brine*

*S, brine*

Figure 5: Schematic diagram of the NF membrane streams.

**Cooling reject**

**Concentrated brine**

**Seawater Feed Blow down**

> Figure 4 shows the hybrid FO-multi effect distillation (FO-MED) system for high TBT MED. In this type of hybrid system, the brine reject from the last effect will be considered as a draw solution stream, while the condenser cooling seawater will be used as a feed solution stream. Permeate water will transport across the FO membrane from the feed to the draw solution side while monovalent and multivalent ions are rejected by the FO membrane. After leaving the FO membrane, the concentrated seawater is dumped back to the sea water. Simultaneously, the diluted draw solution from the FO process is circulated to the MED evaporator as a makeup feed. Inside the MED evaporator, fresh water is extracted from the draw solution by evapo‐ ration and is condensed in the consecutive MED effect. The distilled water is collected at the last effect and is directed to the distilled tank. The brine (un-evaporated) through MED effect is collected in the last effect at a high salinity is directed to the FO again as a draw solution. **Diluted brine Diluted brine Permeate Concentrated brine** Figure 3: Process flow diagram of hybrid FO-MSF process B/H **Rejection FO membrane Seawater Feed Cooling reject Diluted brine Permeate Concentrated brine Cooling reject MSF recovery section**

Last

Figure 4: Process flow diagram of hybrid FO-MED process **Figure 4.** Process flow diagram of hybrid FO-MED process

Condensate return

**MSF recovery section**

#### **3. Methodology**

#### **3.1. Mathematical model development of NF**

*W, feed S, feed W, permeate* Figure 5 illustrates the input and output parameters used for the mass and energy balance equations of the NF membrane [13-14].

FO membrane

Figure 4: Process flow diagram of hybrid FO-MED process

Figure 5: Schematic diagram of the NF membrane streams. **Figure 5.** Schematic diagram of the NF membrane streams

Mass balance is written as follows:

$$W\_{f,f} = W\_{p,f} + W\_{b,f} \tag{1}$$

$$S\_{f,f} = S\_{p,f} + S\_{b,f} \tag{2}$$

The following relation defines the rate of water passage through a semipermeable mem‐ brane [14]:

$$W\_{p,j} = (\Delta P\_j - \sigma \Delta \pi\_j) \times K\_w \times A\_j \times TCF \times FF \times \rho\_{p,j} \tag{3}$$

$$
\Delta P\_j = \overline{P}\_j - P\_{p,j} \tag{4}
$$

$$
\Delta \boldsymbol{\pi}\_{\ j} = \bar{\boldsymbol{\pi}}\_{\ j} - \boldsymbol{\pi}\_{p,j} \tag{5}
$$

$$
\bar{P}\_j = 0.\mathcal{S}(P\_{f,j} + P\_{b,j})\tag{6}
$$

As the seawater salt concentrations ratio is almost constant, an approximation for value in kPa can be given as [13]:

$$
\pi = 6.89\,\text{5} \times \frac{38.5 \times C\_{\text{ $\beta$ NaCl} $} \times (T + 273)}{1000 + C\_{\text{$ \beta $NaCl}$ }} \tag{7}
$$

$$C\_{\rho\_{\text{flv}}\text{Cl}} = 0.934348 \times C\_{\rho\_{\text{fl}}} - 0.54169 \tag{8}$$

The rate of salt flow through the membrane is defined as:

$$S\_{p,j} = (C\_{m,j} - C\_{p,j}) \times K\_s \times A\_j \times TCF + (1 - \sigma) \times J\_{v,j} \times \bar{C} \times K\_s \times A\_j \times TCF \tag{9}$$

$$J\_{v,j} = (\Delta P\_j - \sigma \Delta \pi\_j) \times K\_w \times FF \times TCF \left(\mathfrak{m} / \text{s}\right) \tag{10}$$

Where, the temperature factor correction (TCF) is calculated using the following equations [14]:

$$TCF = e^{8.899 \times \frac{T-2S}{T+2T3}}, \text{ for } T \ge 2S \text{°C} \tag{11}$$

$$TCF = e^{\frac{11.678 \times \frac{T-2S}{T+2\Im S}}{T+2\Im S}}, \text{ for } T \ge 2S \text{°C} \tag{12}$$

$$C\_{p,j} = S\_{p,j} \times \rho\_{p,j} / (S\_{p,j} + W\_{p,j}) \tag{13}$$

A material balance within the mass transfer boundary layer near the membrane wall between the solute carried to the membrane by convection and the solute carried away by diffusion yields an expression that quantifies concentration polarization:

Mass balance is written as follows:

brane [14]:

222 Desalination Updates

can be given as [13]:

*WWW f j pj bj* , ,, = + (1)

*S SS f j pj bj* , ,, = + (2)

r

D=- *P PP j j pj* , (4)

*j j pj* , (5)

*<sup>C</sup>* (7)

, , = + 0.5( ) *P PP <sup>j</sup> f j bj* (6)

*p j* (3)

The following relation defines the rate of water passage through a semipermeable mem‐

, , *W P K A TCF FF pj j j w j* =D -D ´ ´ ´ ´ ´ ( )

\_\_

\_ D= p pp

As the seawater salt concentrations ratio is almost constant, an approximation for value in kPa

´ ´+

*C T*

+ *fbNaCl*

s

 *pj mj pj s j* = - ´ ´ ´ + - ´ ´´ ´ ´ *vj s j S C C K A TCF J C K A TCF* (9)

*fbNaCl*

= ´- 0.934348 0.54169 *C C fbNaCl fb* (8)

*<sup>T</sup> TCF e T* (11)

*<sup>T</sup> TCF e T* (12)

/( ) *C S SW pj pj pj pj pj* (13)

*K F TCF* (*mF s*) (10)

38.5 ( 273) 6.895 1000

s p

\_

p

The rate of salt flow through the membrane is defined as:

= ´

, ,, , ( ) (1 )

s p

*Jvj j j w* , =D -D ´ ´ ´ () / *P*

Where, the temperature factor correction (TCF) is calculated using the following equations [14]:

<sup>25</sup> 8.859 <sup>273</sup> , 5r <sup>2</sup> Cfo - ´ <sup>+</sup> <sup>=</sup> ³ ° *T*

<sup>25</sup> 11.678 <sup>273</sup> , 5r <sup>2</sup> Cfo - ´ <sup>+</sup> <sup>=</sup> ³ ° *T*

, ,, , , =´ + r

$$\varphi = \frac{C\_m - C\_p}{C\_b - C\_p} = e^{J\_w/k} \tag{14}$$

The Umm-Lujj NF-RO plant [20] is considered as a case study to verify the mathematical model of the NF membrane equation (1-14), as well as to estimate the permeate constant Kw and the solute constant Ks.

This plant consists of 27 pressure vessels and six NF elements per vessel. The feed characteristic is 360 m3 /hr, the temperature is 32°C, and the salinity is 45.46 g/l. The applied feed pressure is 25 bars. The data from the Umm-Lujj plant, shown in Table (1), are used as the input data of VDS [21-25] software as shown in Figure 6.

**Figure 6.** VDS interface of the NF system with pressure exchanger

The VDS simulates the Umm-Lujj plant of NF to estimate the permeate production and the exact value of the membrane constants Kw and Ks. After several runs, the membrane water permeability *Kw* of the considered NF membrane is determined as follows:

$$K\_w = 5.8 \times 10^{-9} \text{ m}^3/\text{m}^2 \text{.s.kPa} \tag{15}$$

The membrane salt permeability coefficient *Ks* is estimated as follows:

$$K\_s = 9 \times 10^{-8} \tag{16}$$

Using the estimated values Kw and Ks, the VDS results are compared against the typical plant as shown in Table (1). The comparison results show a good agreement between the VDS results and the typical real plant.


**Table 1.** Comparison between VDS and Umm-Lujj typical results

#### \*Input data to the VDS

The VDS program is used to size the NF system to produce 226 m3 /hr, which represents onethird of the makeup feed, which is required for 1 MIGD MSF. The required number of pressure vessels is calculated as 29 with 174 membrane elements. The calculated system recovery ratio is 65%. The high-pressure pump is assigned by 25 bars. Three units of pressure exchangers are used to recover an electrical energy of 0.07 MW. Each unit's capacity is 44 m3 /hr, and the percentage of salt increase is only 4.6%. The net pumping power required is 0.21 MW, and the specific power consumption is 0.94 kWh/m3 .

#### **3.2. Mathematical model of forward osmosis (FO)**

The VDS simulates the Umm-Lujj plant of NF to estimate the permeate production and the exact value of the membrane constants Kw and Ks. After several runs, the membrane water

Using the estimated values Kw and Ks, the VDS results are compared against the typical plant as shown in Table (1). The comparison results show a good agreement between the VDS results

/hr, \* 360 360 -

/hr 245 234 4.7%

**Variable VSP results Umm-Lujj % Error**

Feed salinity, TDS, g/l, \* 45.46 45.46 - Stages No.\* 1 1 - No. of pressure vessels, \* 27 27 - Feed temperature, °C, \* 32 32 - Fouling factor, \* 0.95 NA - Feed pressure, bar, \* 25 25 - Elements No. per vessel, \* 6 6 -

Recovery ratio 0.68 0.65 4.6% Permeate salinity, TDS, mg/l 29.11 28.26 3%

third of the makeup feed, which is required for 1 MIGD MSF. The required number of pressure vessels is calculated as 29 with 174 membrane elements. The calculated system recovery ratio is 65%. The high-pressure pump is assigned by 25 bars. Three units of pressure exchangers are

percentage of salt increase is only 4.6%. The net pumping power required is 0.21 MW, and the

.

used to recover an electrical energy of 0.07 MW. Each unit's capacity is 44 m3

**Table 1.** Comparison between VDS and Umm-Lujj typical results

specific power consumption is 0.94 kWh/m3

The VDS program is used to size the NF system to produce 226 m3

93 2 5. m 8 10 .s.kPa m / - *Kw* = ´ (15)

<sup>8</sup> 9 10- *Ks* = ´ (16)

/hr, which represents one-

/hr, and the

permeability *Kw* of the considered NF membrane is determined as follows:

The membrane salt permeability coefficient *Ks* is estimated as follows:

and the typical real plant.

Feed flow rate, m3

224 Desalination Updates

Permeate flow rate, m3

\*Input data to the VDS

Forward osmosis is the transport of water across a selectively permeable membrane from a region of higher water chemical potential (feed solution) to a region of lower water chemical potential (draw solution). Consequently, a less concentrated draw solution is being produced, which may be further treated to extract freshwater. Obviously, there are two key problems that must be solved to make the technology go out of the laboratory. One is continued improvement and optimization of the selectively permeable membrane, which allows passage of water but rejects most solute molecules or ions. The other is the identification of optimal osmotic agents and its corresponding recovery processes for the supply of the osmosis pressure difference, which is the driving force of the FO process. However, the FO membrane water flux is far lower than the anticipated when the membrane used is asymmetric. The primary reason for this finding is the fact that both FO are accompanied by internal concentration polarization (ICP) as shown in Figure 7.

Figure 6: VDS Interface of NF system with pressure exchanger

Figure 7: Concentration profile in asymmetrical FO membrane **Figure 7.** Concentration profile in asymmetrical FO membrane

Concentrated Draw solution Diluted Draw solution Sea water Feed Concentrated Sea water Feed Permeate FO membrane Figure 8: Schematic of FO membrane process. In the work of Jung et al. [26], FO performance (permeate flux and recovery rate) of a 10 cm ×10 cm plate and frame type membrane is investigated via a numerical simulation based on the mass conservation theorem. The FO membrane orientation, flow direction of feed and draw solutions, flow rate, and solute resistivity (K) are simulated. The case of draw solution facing the active layer displays a relatively higher performance than the feed solution facing the active layer [26]. The numerical results showed that the membrane performance is much more sensitive to the physical membrane property parameter rather than the flow rate and flow direction. However, the simulation methodology does not consider fouling and reverse solute diffusion. Also, the performance of the FO in a relatively large size needs to be explored to come up with a concrete recommendation to the commercialization phase.

A simple schematic of the FO process is shown in Figure 8. The mathematical model of FO membrane is developed as shown in equations (17-26). By knowing the specifications of the FO membrane,, and the membrane area, the outlet stream can be calculated.

FO membrane

Figure 7: Concentration profile in asymmetrical FO membrane

Figure 6: VDS Interface of NF system with pressure exchanger

Porous

Active layer

Feed Solution

Low

concentration

**Figure 8.** Schematic of the FO membrane process

Draw

> High

concentration

Solution

$$W\_{DS, \text{outlet}} - W\_{DS, \text{inlet}} = W\_p \tag{17}$$

$$S\_{DS,out} - S\_{DS,out} = S\_p \tag{18}$$

$$W\_{F\ $,in} - W\_{F\$ ,out} = W\_p \tag{19}$$

$$\rm{S}S\_{FS,m} - S\_{FS,out} = S\_p \tag{20}$$

$$W\_p = J\_w \times A \times \rho \tag{21}$$

$$S\_p = J\_s \times A \tag{22}$$

The general flux equation in the FO process is

$$J\_w = A \left(\pi\_D - \pi\_F\right) \tag{23}$$

where *A* is the water permeability coefficient, and and denote the osmotic pressures of the draw and feed solution, respectively. The osmotic pressure can be determined by the modified Van't Hoff equation as:

$$
\Delta \pi = \frac{N\_{\text{low}} R\_g T \Delta C}{M\_w} \tag{24}
$$

where T, and indicate the ionization number of the solution, the ideal gas constant, the absolute temperature, the salt concentration difference of the solution across the membrane, and the molecular weight of the salt, respectively. For asymmetrical membrane, internal ICP occurs within the porous support as shown in Figure 7. However, when the reverse solute is consid‐ ered [27], and for the case of the draw solution facing the porous layer, the water flux and solute flux are:

$$J\_w = A \left( \frac{\pi\_o \exp\left(-\frac{J\_w S}{D}\right) - \pi\_\nu \exp\left(\frac{J\_w}{k}\right)}{1 + \frac{B}{J\_w} \left[ \exp\left(\frac{J\_w}{k}\right) - \exp\left(-\frac{J\_w S}{D}\right) \right]} \right) \tag{25}$$

$$J\_x = B \left[ \frac{C\_B \exp\left(-\frac{J\_w S}{D}\right) - C\_\prime \exp\left(\frac{J\_w}{k}\right)}{1 + \frac{B}{J\_w} \left[ \exp\left(\frac{J\_w}{k}\right) - \exp\left(-\frac{J\_w S}{D}\right) \right]} \right] \tag{26}$$

where B is the salt permeability coefficient.

Figure 6: VDS Interface of NF system with pressure exchanger

Porous

Figure 7: Concentration profile in asymmetrical FO membrane

Sea water Feed Concentrated Sea water Feed

Concentrated Draw solution Diluted Draw solution

Permeate

FO membrane

Figure 8: Schematic of FO membrane process.

*WJA p w* = ´´

*J A w DF* = - (p p

p

1

1

*w*

*s*

p

*w*

*w*

D = *ion g*

where *A* is the water permeability coefficient, and and denote the osmotic pressures of the draw and feed solution, respectively. The osmotic pressure can be determined by the modified

D

*w*

where T, and indicate the ionization number of the solution, the ideal gas constant, the absolute temperature, the salt concentration difference of the solution across the membrane, and the molecular weight of the salt, respectively. For asymmetrical membrane, internal ICP occurs within the porous support as shown in Figure 7. However, when the reverse solute is consid‐ ered [27], and for the case of the draw solution facing the porous layer, the water flux and

> pæ ö æ ö æö ç ÷ ç ÷ ç÷ - -

*JS J*

*w w*

*w w*

*w w*

*w w*

è ø èø <sup>=</sup> é ù æö æ ö + -- ê ú ç÷ ç ÷ è ø ë û èø è ø

*B J JS exp exp Jk D*

æ ö æ ö æö ç ÷ ç ÷ ç÷ - -

*B J JS exp exp Jk D*

è ø èø <sup>=</sup> é ù æö æ ö + -- ê ú ç÷ ç ÷ è ø ë û èø è ø

*D F*

*JS J C exp C exp D k J B*

*D F*

*exp exp D k J A*

*N RT C*

r

Draw

**Figure 8.** Schematic of the FO membrane process

226 Desalination Updates

The general flux equation in the FO process is

Van't Hoff equation as:

solute flux are:

High

concentration

Solution

Active layer

Feed Solution

, , *W WW DS outlet DS inlet p* - = (17)

, , *SS S DS out DS inlet p* - = (18)

*WW W FS in FS out p* , , - = (19)

*SS S FS in FS out p* , , - = (20)

*S JA p s* = ´ (22)

(21)

) (23)

*<sup>M</sup>* (24)

(25)

(26)

Low

concentration

The water flux and reverse solute equations are difficult to solve analytically because these equations are dependent on water flux and solute passage including in concentration polari‐ zation terms. Thus, using a program to solve these equations will be presented.

The FO is still facing many challenges such as internal concentration polarization, which requires breakthrough in the molecular design of high-performance FO membrane. On the other hand, development of draw solution with low cost and low energy consumption required for recovery is urgently needed. To mitigate ICP, the FO membranes must have characteristics of high permeability and hydrophobicity with a small structure parameter, while the preferred draw solutes must have diffusion coefficient, reasonable molecular size, and low viscosity [52].

The VDS Software interface is developed to consider a case study as shown in Figure 9. As shown in Figure 9, the specified feed solution of (1680 t/h and 45 g/l) is pumped to the FO membrane against a draw solution (NaCl) (1200 t/h and 90 g/l). Three elements per vessel are arranged. Two pressure vessels are placed in parallel. The VDS simulated the case, and the results are presented in the same interface as shown in Figure 9. The diluted flow rate is calculated as 1744 t/h and 64.34 g/l, while the concentrated feed solution is calculated as 1135.8 t/h and 66.46 g/l. Based on these calculations, the system recovery ratio is calculated as 45%, and the specific power consumption is calculated as 0.55 kWh/m3 . So far, the effect of concen‐ tration polarization is not considered yet.

**Figure 9.** Interface of VDS software for FO process
