**4. Technoeconomic analysis**

#### **4.1. NF-MSF**

The first MSF of 0.5 MGD per unit evaporator was built in 1957 in Kuwait using the once through MSF-OT configuration by the Westinghouse Company [28]. The design was modified according to the recommendation of the client, the Ministry of Electricity and Water, in Kuwait and of engineers for reliable operation. For some time, the market was dominated by the once through (MSF-OT) due its simplicity and high thermodynamic efficiency. However, due to high oxygen and CO2 gas liberation in addition to large amount of feed water to be pre-treated, the market was forced to shift to the brine recirculation configuration (MSF-BR). The first 19 stage 1 MGD MSF-BR plant was built by Weir Company in 1959 in Kuwait [28]. The developed specifications led to more reliable, easy-to-operate-and-maintain, and longer life units. Now, the MSF evaporator production capacity was increased dramatically through the years to reach 20 MGD in UAE, and designs of 25-30 MIGD are available. The disadvantage of the MSF-BR system is the higher brine concentration, which increases the potential for having scale deposits on the heat transfer surfaces and for the liquid boiling point elevation (BPE), thus penalizing the coefficient of heat transfer and the available condensing temperature difference, respec‐ tively.

Increasing the MSF unit production (for both new designs and operating units) can be carried out either by: i) increasing the re-circulating brine flow rate, or ii) increasing the flashing range. Increasing the re-circulating brine flow rate is limited, however, by the available pump capacity and the chamber load (flashing brine flow velocity). Increasing the flashing range (TBT–BBT) can be carried out by increasing the top brine temperature (TBT), with hard-scale solution, or reducing the bottom brine temperature (BBT), with lower heat sink temperature (naturally in fall/winter/spring or utilizing deep intake or cooling towers). Increasing TBT is the parameter addressed in this paper.

At high TBT, scale deposits of high seawater brine concentration present a real problem in MSF plants, as they directly affect the heat transfer rates on the heating surface. The main scaleforming constituents are calcium (Ca++), magnesium (Mg++), bicarbonate (HCO3- ), and sulphate (SO4--) ions. On heating, bicarbonate decomposes into carbonate CO3- , which reacts with Ca+ + forming calcium carbonate (CaCO3) that precipitates on the heat transfer surface (if saturation limits are exceeded). At high temperature, magnesium hydroxide (MgOH) will also be formed. At higher temperature of >120°C, non-alkaline calcium sulphate (CaSO4) precipitates if saturation limits are also exceeded, due to inverse solubility. Formation of alkaline scale (CaCO3 and MgOH) can be controlled by lowering the pH (acid additives) or by anti-scalant. Non-alkaline (hard) scale (such as CaSO4) is only controlled nowadays by limiting the TBT below 120°C.

Scale deposits have a direct influence on the thermal units' performance and water cost. Sulphate scales are a result of the direct crystallization of anhydrite (CaSO4), hemihydrate (CaSO4 0.5H2O), or gypsum (CaSO4 2H2O) from seawater once their solubility limits are exceeded as shown Fig. 10 [29]. Most of the deposited calcium sulphate found in seawater desalination plants is in the form of hemi-hydrate. The sulphate minerals are insoluble in common chemicals, and their development inside a distiller should be avoided by all means.

**Figure 10.** Phase diagram of CaSO4

**4. Technoeconomic analysis**

The first MSF of 0.5 MGD per unit evaporator was built in 1957 in Kuwait using the once through MSF-OT configuration by the Westinghouse Company [28]. The design was modified according to the recommendation of the client, the Ministry of Electricity and Water, in Kuwait and of engineers for reliable operation. For some time, the market was dominated by the once through (MSF-OT) due its simplicity and high thermodynamic efficiency. However, due to high oxygen and CO2 gas liberation in addition to large amount of feed water to be pre-treated, the market was forced to shift to the brine recirculation configuration (MSF-BR). The first 19 stage 1 MGD MSF-BR plant was built by Weir Company in 1959 in Kuwait [28]. The developed specifications led to more reliable, easy-to-operate-and-maintain, and longer life units. Now, the MSF evaporator production capacity was increased dramatically through the years to reach 20 MGD in UAE, and designs of 25-30 MIGD are available. The disadvantage of the MSF-BR system is the higher brine concentration, which increases the potential for having scale deposits on the heat transfer surfaces and for the liquid boiling point elevation (BPE), thus penalizing the coefficient of heat transfer and the available condensing temperature difference, respec‐

Increasing the MSF unit production (for both new designs and operating units) can be carried out either by: i) increasing the re-circulating brine flow rate, or ii) increasing the flashing range. Increasing the re-circulating brine flow rate is limited, however, by the available pump capacity and the chamber load (flashing brine flow velocity). Increasing the flashing range (TBT–BBT) can be carried out by increasing the top brine temperature (TBT), with hard-scale solution, or reducing the bottom brine temperature (BBT), with lower heat sink temperature (naturally in fall/winter/spring or utilizing deep intake or cooling towers). Increasing TBT is the parameter

At high TBT, scale deposits of high seawater brine concentration present a real problem in MSF plants, as they directly affect the heat transfer rates on the heating surface. The main scale-

 forming calcium carbonate (CaCO3) that precipitates on the heat transfer surface (if saturation limits are exceeded). At high temperature, magnesium hydroxide (MgOH) will also be formed. At higher temperature of >120°C, non-alkaline calcium sulphate (CaSO4) precipitates if saturation limits are also exceeded, due to inverse solubility. Formation of alkaline scale (CaCO3 and MgOH) can be controlled by lowering the pH (acid additives) or by anti-scalant. Non-alkaline (hard) scale (such as CaSO4) is only controlled nowadays by limiting the TBT

Scale deposits have a direct influence on the thermal units' performance and water cost. Sulphate scales are a result of the direct crystallization of anhydrite (CaSO4), hemihydrate (CaSO4 0.5H2O), or gypsum (CaSO4 2H2O) from seawater once their solubility limits are exceeded as shown Fig. 10 [29]. Most of the deposited calcium sulphate found in seawater

), and sulphate

, which reacts with Ca+

forming constituents are calcium (Ca++), magnesium (Mg++), bicarbonate (HCO3-

(SO4--) ions. On heating, bicarbonate decomposes into carbonate CO3-

**4.1. NF-MSF**

228 Desalination Updates

tively.

+

below 120°C.

addressed in this paper.

Increasing the TBT with hard-scale solution can be carried out by: i) introducing hightemperature anti-scalant, and ii) reducing hard-scale ions to avoid it from reaching the saturation conditions. The first is not yet available through the use of nano-filtration (NF) membrane system for make-up feed water pretreatment.

The application of NF in seawater desalination has gained significant attention in the desali‐ nation industry due to the selective removal of divalent ions. The SWCC R&D team [30-32] carried out extensive experiments on an MSF test pilot unit with NF as the pretreatment. NF pressure was 24 bars, and its recovery ratio ranged from 60% to 65%. The total concentration of the sulphate and calcium ions of the brine recycle was at a TBT of 130°C, and the makeup entirely formed from NF permeate was below the solubility limits. This result indicated the possibility of operating the MSF plant safely and without any scaling problem at TBT equal to or higher than 130°C. However, many questions on the adding of capital cost which might result in saving in operational cost still need clear answers.

The NF was originally applied to reject electrolytes and obtain ultra-pure water with high volume flux at low operating pressure, as most membranes have either positive or negative charge due to their compositions [33]. The NF membrane possesses a molecular weight cutoff of about hundreds to a few thousands, which is intermediate between reverse osmosis (RO) membranes and ultra-filtration (UF). The pore radii and fixed charge density of practical membranes were evaluated from permeation experiments of different neutral solutes of sodium chloride. The pore radii of these NF membranes were estimated to range from 0.4 to 0.8 nm [33].

The flexible and powerful tool ''Visual Design and Simulation program (VDS)'' is used to perform process and techno-economical calculations. VDS was developed for the design and simulation of different types and configurations of the desalination processes [21-25]. Typical desalination processes are simulated to show the wide scope and high capability of the developed package. The description of the VDS software and discussions on how to access and handle the package are presented in [21-25]. In this work, the scope of the VDS program will be extended to develop and build up an NF system and a new MSF configuration model. The NF system's mathematical model will be verified using typical NF-RO plant data.

Table (2) shows the CAPEX cost analysis of the NF system, which produces 226 m3 /hr. The direct costs of the purchased equipment (membrane section, filters, pumps, valves, and piping) are included. The indirect costs of the building structure, engineering, and project develop‐ ment are also included. The intake cost is not included and is assumed to be included in the MSF CAPEX cost. The levelized cost is calculated (based on the 7% interest rate and 15 year life span) as 0.0775 \$/m3 of the NF permeate as shown in column three of Table (2).


**Table 2.** CAPEX cost analysis of the NF system [14]

Table (3) shows the operational cost of the NF system, which includes labor, O&M, NF membrane replacement, electricity, and chemicals. The analysis showed that the cost of electricity represents the biggest chunk of the total OPEX, and the specific operational cost is 0.0566 \$/m3 of the NF permeate. From both Table (2) and Table (3), the calculated unit permeate cost is 0.134 \$/m3 .

Figure 11 shows the interface of the existing 5,000 m3 /day MSF-BR desalination plant at TBT=110°C [14]. The evaporator consists of 20 stages – 17 for the heat recovery section and 3 stages for the heat rejection section. The extracted steam from the power side is directed to the brine heater as a heat source. Sea water flows through the tubes of the heat rejection section condensers as a coolant. Part of this coolant outlet is used as a make-up, and the remaining coolant is rejected back to the sea. The make-up is directed to the de-aerator, and pretreatment chemicals are added, then mixed with a portion of the last-stage brine. The circulation pump circulates the diluted mixed brine to the condensers of the heat recovery section. The tube


**Table 3.** OPEX cost analysis of the NF system [14]

Table (2) shows the CAPEX cost analysis of the NF system, which produces 226 m3

**Items Cost, \$ Specific cost, \$/m3**

PV, pass 1 26,100.00 0.001510285 Element, NF 87,000.00 0.005034285 Pumps 118,026.20 0.006829626 PX/turbine 115,359.61 0.006675323 Piping and valves 188,939.88 0.010933071 Filters 249,015.74 0.014409381 Others, building, start-up 258,803.85 0.014975773 Subtotal 1,043,245.28 0.060367744

Engineering design 131,455.62 0.007606724 Financial 164,319.52 0.009508405 Sub-total 295,775.14 0.017115129 TCI 1,339,020.42 0.077482874

Table (3) shows the operational cost of the NF system, which includes labor, O&M, NF membrane replacement, electricity, and chemicals. The analysis showed that the cost of electricity represents the biggest chunk of the total OPEX, and the specific operational cost is

TBT=110°C [14]. The evaporator consists of 20 stages – 17 for the heat recovery section and 3 stages for the heat rejection section. The extracted steam from the power side is directed to the brine heater as a heat source. Sea water flows through the tubes of the heat rejection section condensers as a coolant. Part of this coolant outlet is used as a make-up, and the remaining coolant is rejected back to the sea. The make-up is directed to the de-aerator, and pretreatment chemicals are added, then mixed with a portion of the last-stage brine. The circulation pump circulates the diluted mixed brine to the condensers of the heat recovery section. The tube

of the NF permeate. From both Table (2) and Table (3), the calculated unit permeate

/day MSF-BR desalination plant at

life span) as 0.0775 \$/m3

Direct cost

230 Desalination Updates

Indirect cost

0.0566 \$/m3

cost is 0.134 \$/m3

**Table 2.** CAPEX cost analysis of the NF system [14]

.

Figure 11 shows the interface of the existing 5,000 m3

direct costs of the purchased equipment (membrane section, filters, pumps, valves, and piping) are included. The indirect costs of the building structure, engineering, and project develop‐ ment are also included. The intake cost is not included and is assumed to be included in the MSF CAPEX cost. The levelized cost is calculated (based on the 7% interest rate and 15 year

of the NF permeate as shown in column three of Table (2).

/hr. The

materials used in this plant are CuNi 90/10 for the brine heater and heat recovery section and CuNi 70/30 for the heat rejection section. The evaporator length is 29 m, the width is 7 m, and the height is 2.5 m. The design conditions are 27°C for seawater, and the brine velocity inside the tube is 2 m/s. The working pressure of the de-aerator is 0.055 bars, which is lower than the make-up saturation temperature of 38°C.

**Figure 11.** Interface of the existing MSF-BR desalination plant at TBT=110 C [14]

Figure 12 shows the interface the MSF-BR with the NF system, which allows increasing the TBT to 130°C. The NF system treats one-third of the make-up. The feed of the NF system is

**Figure 12.** Interface of the NF-MSF-BR desalination plant [14]

extracted from the cooling reject stream (48 g/l) as shown in Figure (5). The NF permeate is mixed with the remaining make-up and directed to the de-aerator. The mixed make-up of low salinity of 43 g/l (15 % les) flows to the last stage of the heat rejection section. Due to the increase in the TBT from 110°C to 130°C, the distillate production increases by 19%. There is no increase of the GOR, as the heating steam increased also by 19%.

Table (4) shows that the CAPEX of the NF-MSF-BR system is 65.5% higher than that for the conventional MSF. Table (5) shows that the operating cost of the NF-MSF-BR system also increased by 22.4% higher than that for the conventional MSF.


**Table 4.** CAPEX analysis of MSF and NF-MSF-BR [14]


**Table 5.** OPEX analysis of MSF and NF-MSF-BR [14]

extracted from the cooling reject stream (48 g/l) as shown in Figure (5). The NF permeate is mixed with the remaining make-up and directed to the de-aerator. The mixed make-up of low salinity of 43 g/l (15 % les) flows to the last stage of the heat rejection section. Due to the increase in the TBT from 110°C to 130°C, the distillate production increases by 19%. There is no increase

Table (4) shows that the CAPEX of the NF-MSF-BR system is 65.5% higher than that for the conventional MSF. Table (5) shows that the operating cost of the NF-MSF-BR system also

**CAPEX, 1 MIGD Conventional NF-MSF % diff**

Evaporator 1,040,551.34 1,040,551.344 - Pumps 306,223.35 306,223.350 - Pipes, valves, I&C 302666.77 302,666.770 - Intake 394560 394,560.000 -

Total 2,044,001.46 3,383,021.88 65.5%

of the GOR, as the heating steam increased also by 19%.

**Figure 12.** Interface of the NF-MSF-BR desalination plant [14]

232 Desalination Updates

increased by 22.4% higher than that for the conventional MSF.

Item COST, US\$ COST, US\$

NF system - 1,339,020.42

**Table 4.** CAPEX analysis of MSF and NF-MSF-BR [14]

Table (6) shows that the levelized CAPEX cost of MSF-BR at TBT=130°C is 16% lower than that for the conventional MSF at TBT=110°C. This is due to the increase in distillate production of 19%. Also, due to the increase in the productivity, the specific OPEX reduced by 2.5%. Howev‐ er, due to adding the NF system, the levelized OPEX of NF-MSF at TBT=130°C is 2.65% higher, while the specific CAPEX of NF-MSF is 28.7% higher than that for the conventional MSF.

The unit product cost of NF-MSF is 5.4% higher than that of the MSF plant. The analysis of these CAPEX and OPEX results shows that the OPEX cost has significant effect on the total unit water cost. This can be concluded that adding the NF system to an existing MSF plant (just to increase the production) is not enough to reduce the unit product cost.


**Table 6.** Levelized cost of MSF and NF-MSF-BR [14]

A modified MSF-DM configuration has been proposed as shown in Figure 13. In this MSF-DM configuration, the heat rejection section is removed, and the bottom part of the de-aerator is utilized as a mixer where part of the last stage brine is mixed with de-aerated make-up. The new configuration is half-way between brine recirculation MSF-BR and once through MSF-OT and will benefit from both techniques and overcome the limitation encountered through operation. The GOR of the MSF-DM configuration at TBT=110°C could be as high as 12.

The MSF-DM design configuration is targeting high MSF GOR to be adopted in solar energy applications (high GOR is also needed, as the cost of energy is increasing). As the capital cost in solar energy systems is expensive, it will be cost-effective to develop high-performance MSF to reduce the CAPEX of the solar energy systems. A high-performance MSF system requires a combination of more evaporating stages, and more heat transfer surface area sequentially increases the MSF CAPEX. The increase in the MSF CAPEX could be balanced by reducing the MSF OPEX, and accordingly, CAPEX reduction of the solar energy system will be the main contribution to the developed system. **Figure 12: Interface NF-MSF-BR Desalination plant, [14].**

**Figure 13: The interface of the new MSF-DM for desalination plant, [14]. Figure 13.** The interface of the new MSF-DM for desalination plant [14]

Figure 14 shows the configuration of NF with the newly developed de-aeration brine mix NF-MSF-DM system to reduce the operational cost (OPEX). NF enables increasing the TBT to 130°C, while the MSF-DM enables increasing the GOR. As shown in Figure 14, the stage number of MSF-DM increased to 35, which is 75% higher than that of the conventional MSF-BR. The 61.5% of the last-stage brine is mixed with the de-aerated make-up flow of 675 m3 /hr. The make-up is diluted from 48 g/l to 43 g/l using NF system permeate of TDS=28 g/l. This mixed is directed to the MSF condensers at 32.8°C, which is 15% lower than that for the conventional MSF (38°C). This lower temperature of coolant enhances the heat transfer process (condensation). However, the reducing cooling water reduces the LMTD across condenser compared with that of the conventional one. This explains why the heating surface area of MSF-DM was increased by 72%. One feature of increasing the heat transfer area of the heat

**Figure 14: NF-MSF-DM configuration**

Techno-Economics of Hybrid NF/FO with Thermal Desalination Plants http://dx.doi.org/10.5772/60207 235

**Figure 14.** NF-MSF-DM configuration

A modified MSF-DM configuration has been proposed as shown in Figure 13. In this MSF-DM configuration, the heat rejection section is removed, and the bottom part of the de-aerator is utilized as a mixer where part of the last stage brine is mixed with de-aerated make-up. The new configuration is half-way between brine recirculation MSF-BR and once through MSF-OT and will benefit from both techniques and overcome the limitation encountered through operation. The GOR of the MSF-DM configuration at TBT=110°C could be as high as 12.

The MSF-DM design configuration is targeting high MSF GOR to be adopted in solar energy applications (high GOR is also needed, as the cost of energy is increasing). As the capital cost in solar energy systems is expensive, it will be cost-effective to develop high-performance MSF to reduce the CAPEX of the solar energy systems. A high-performance MSF system requires a combination of more evaporating stages, and more heat transfer surface area sequentially increases the MSF CAPEX. The increase in the MSF CAPEX could be balanced by reducing the MSF OPEX, and accordingly, CAPEX reduction of the solar energy system will be the main

**Figure 12: Interface NF-MSF-BR Desalination plant, [14].**

E-225

**Figure 13: The interface of the new MSF-DM for desalination plant, [14].**

Figure 14 shows the configuration of NF with the newly developed de-aeration brine mix NF-MSF-DM system to reduce the operational cost (OPEX). NF enables increasing the TBT to 130°C, while the MSF-DM enables increasing the GOR. As shown in Figure 14, the stage number of MSF-DM increased to 35, which is 75% higher than that of the conventional MSF-BR. The 61.5% of the last-stage brine is mixed with the de-aerated make-up flow of 675 m3

The make-up is diluted from 48 g/l to 43 g/l using NF system permeate of TDS=28 g/l. This mixed is directed to the MSF condensers at 32.8°C, which is 15% lower than that for the conventional MSF (38°C). This lower temperature of coolant enhances the heat transfer process (condensation). However, the reducing cooling water reduces the LMTD across condenser compared with that of the conventional one. This explains why the heating surface area of MSF-DM was increased by 72%. One feature of increasing the heat transfer area of the heat

**Figure 14: NF-MSF-DM configuration**

E-224

E-226

To stage 20

Vacuum deaerator

E-223

Condensate hotwell E-229

E-230

Sea water

Distillate tank

Feed

Blow down

/hr.

contribution to the developed system.

Motive steam

V-14

B/H

E-220

Steam from power plant

234 Desalination Updates

Condensate

recovery section

1 17

E-221 E-231

E-228 E-227

2 8 14

**Figure 13.** The interface of the new MSF-DM for desalination plant [14]

recovery section is reducing the temperature difference across the brine heater, which sequentially increases the brine heater surface area. Increasing the heat transfer area of the heat recovery section increases the recovered energy, thus minimizing the external source of heating. Reducing the source of heating (steam) for fixed capacity will increase the GOR. The process calculations show that the GOR is 100% higher than that of MSF-BR (see Table (7)). This means that steam consumption is reduced by 100 %.

Table (7) shows the process calculation of MSF-DM at TBT=130°C compared with the conven‐ tional MSF at TBT=110°C. The GOR of MSF-DM-NF is twice of the conventional MSF; however, the heat transfer area is increased by 72%. Table (7) and Figure 14 show that the intake seawater of MSF-DM is 42% lower than that of MSF-BR. This, in turn, would reduce the seawater supply pump capacity, as well as the intake civil work. One feature of MSF-DM, the make is the same value of the conventional, which leads to having the same chemical cost of treatment and same manufacturing cost of de-aeration.

Table (7) shows that the specific power consumption of MSF-DM is 27% higher than that of MSF-BR. This is because of the increase of the friction loss due to the increase of the stage number by 75%. The evaporator length is increased by 142% in the case of MSF-DM; the evaporator width is decreased by 2%, while height is increased by 7% as shown in Table (7).

The purchased equipment cost (PEC) of these components is estimated based on recent market prices. In cases when data about the real installation cost of the desalination plant are scarce, the PEC of the individual components could be calculated based on cost relations. These relations of estimating the capital and operating costs of the components, such as pumps, valves, piping, and instrumentations are presented in [14].


**Table 7.** Process calculation of MSF and MSF-DM [14]

A detailed cost breakdown is shown in Table (8). The evaporator (shell and tubes, de-aerator) cost of MSF-DM is 47% higher due to the increase in the heat surface area by 75% as shown in Table (7). Evaporator manufacturing cost, including the labor cost, of MSF-DM is 52% higher than that of MSF-BR. The costs of pumps, piping, valves, and I&C control of MSF-DM are lower than that of the conventional system due to the removal of the heat rejection section. The cost analysis shows that the intake construction cost of MSF-DM is 42% lower than that of the conventional one due to lower seawater flow rate. So the increase of MSF-DM evaporator cost is partially compensated by the cost reduction in auxiliaries and intake cost. The total capital cost (CAPEX) of the proposed configuration, MSF-DM, is 6% higher than that of the conven‐ tional MSF-BR. However, the total CAPEX cost of the NF-MSF-DM system is 71% higher than that of the conventional MSF. The main increase in CAPEX is contributed to the additional NF system.

Table (9) shows the OPEX items for both the conventional MSF-BR and the MSF-DM config‐ urations. The cost of the steam and electricity is calculated based on an average of 80 \$/barrel oil price and the recent purchase cost of power generation cycle [14]. The cost of the lowpressure steam is directed to the desalination plant, and the steam utilized for power genera‐ tion is allocated based on exergy analysis [14]. Using levelization method through 20 years and 7%, the specific cost of low pressure steam is calculated as 7.5 \$/m3 of steam, and the cost of the generated electricity is 0.043 \$/kWh.


**Table 8.** CAPEX analysis of MSF and NF-MSF-DM configurations [14]


**Table 9.** OPEX analysis of MSF and MSF-DM-NF [14]

A detailed cost breakdown is shown in Table (8). The evaporator (shell and tubes, de-aerator) cost of MSF-DM is 47% higher due to the increase in the heat surface area by 75% as shown in Table (7). Evaporator manufacturing cost, including the labor cost, of MSF-DM is 52% higher than that of MSF-BR. The costs of pumps, piping, valves, and I&C control of MSF-DM are lower than that of the conventional system due to the removal of the heat rejection section. The cost analysis shows that the intake construction cost of MSF-DM is 42% lower than that of the conventional one due to lower seawater flow rate. So the increase of MSF-DM evaporator cost is partially compensated by the cost reduction in auxiliaries and intake cost. The total capital cost (CAPEX) of the proposed configuration, MSF-DM, is 6% higher than that of the conven‐ tional MSF-BR. However, the total CAPEX cost of the NF-MSF-DM system is 71% higher than that of the conventional MSF. The main increase in CAPEX is contributed to the additional NF

Ton/hr 208 208 0% TBT, C 110 130 18 %

Sea salinity, g/l 48.60 48.6 0% Recycle ratio 0.72 0.615 -15% Recycle salinity, g/l 62.90 60 -5% Blow down salinity, g/l 70.00 70 0% No. of stages 20.00 35 75% Heat transfer area, m2 9,868 16,940 72% Tube length, m 7.45 8 7% Tube diameter, m 19.05 17 -11% GOR 8 16 100% Velocity, m/s 1.98 1.91 -4% SPC, kWh/m3 2.70 3.42 27% Evaporator length, m 29.30 70.9 142% Evaporator height 2.50 2.44 -2% Evaporator width, m 7.45 8 7%

/hr 1370 797 -42 %

/hr 660 675 2%

**MSF-BR NF-MSF-DM % diff**

Table (9) shows the OPEX items for both the conventional MSF-BR and the MSF-DM config‐ urations. The cost of the steam and electricity is calculated based on an average of 80 \$/barrel oil price and the recent purchase cost of power generation cycle [14]. The cost of the lowpressure steam is directed to the desalination plant, and the steam utilized for power genera‐ tion is allocated based on exergy analysis [14]. Using levelization method through 20 years and 7%, the specific cost of low pressure steam is calculated as 7.5 \$/m3 of steam, and the cost

system.

Sea water flow rate, m3

Make-up, m3

236 Desalination Updates

of the generated electricity is 0.043 \$/kWh.

**Table 7.** Process calculation of MSF and MSF-DM [14]

The OPEX cost analysis (Table (9)) shows that the cost for NF-MSF-DM is 31% lower than that for the conventional MSF. The reduction in OPEX contributed to the reduction of the heating steam cost due to higher GOR. The levelized cost showed that the unit product cost of NF-MSF-DM is 21% lower than that of the conventional MSF as shown in Table (10).

As shown in Table (9), the low-pressure steam cost for the MSF-DM configuration is 48% lower than that for the conventional MSF-BR, and the steam consumption of MSF-DM is 100% lower than that consumed by the conventional MSF-BR. This is mainly due to the different steam cost invoked from the power side, as the heat steam temperature is higher in MSF-DM (TBT=130). The electricity cost of the MSF-DM is 27% higher than that of the conventional MSF-BR due to higher pumping power of the same order. The chemical cost is only 2% higher than that for the conventional MSF-BR. This is mainly due to the increase of make-up to be treated. The total number of OPEX items in the proposed configuration, MSF-DM, is 33% lower than that in the conventional MSF-BR, mainly due to low amount of steam consumption.

The annual investment cost (fixed capital cost depreciation rate (FCCDR) per year) of each component in the desalination plant is calculated according to the following relation:

$$\text{Annual Investment} = \text{CAPEX} \times \frac{i \times (1 + i)^n}{(1 + i)^n - 1} \tag{27}$$

Using an interest rate, i, of 7% and number of amortization years, n, of 20 years: then, the operation and maintenance cost is calculated by multiplying the equipment purchase cost by a factor of the equipment cost index. The hourly cost (\$/hr) of the desalination plant is calculated as follows:

$$(hourly - CAPEX = \frac{\text{Total annual investment}}{365 \times 24 \times 0.9} \tag{28}$$

Similarly, the hourly OPEX is calculated as follows:

$$\text{bourly} - \text{OPEX} = \text{LP steam} + \text{Electricity} + \text{Chemicals} \tag{29}$$

Then the unit product cost of the desalted water is calculated as follows:

$$\text{Unit product cost, } \\$/\text{m}^3 = \frac{hourly\\_CAPEX + hourly\\_OPEX}{hourly\\_Product} \tag{30}$$

The levelized cost of capital purchased components and operating invested (chemicals, steam, electricity, O&M) to produce water is calculated as shown in Table (10). The specific OPEX of the MSF-DM is 34% lower than that of the conventional MSF-BR. The specific CAPEX of the MSF-DM is 12% higher than that of the conventional MAF-BR. However, the sum of the total cost invested using the MSF-DM is 34% lower than that of the conventional MSF-BR. Due to adding of the NF system, the specific OPEX of NF-MSF-DM is 20% lower that of the conven‐ tional one, while the specific CAPEX increased by 84% as shown Table (10). The total unit product cost of NF-MSF-DM is 9.5% lower than that of the conventional MSF-BR.


**Table 10.** Levelized cost of MSF and NF-MSF-DM configurations [14]

#### **4.2. Hybrid FO-MSF**

Using an interest rate, i, of 7% and number of amortization years, n, of 20 years: then, the operation and maintenance cost is calculated by multiplying the equipment purchase cost by a factor of the equipment cost index. The hourly cost (\$/hr) of the desalination plant is

Total annual investment

365 24 0.9 - = ´ ´ *hourly CAPEX* (28)

*hourly OPEX* - = LP steam +Electricity+Chemiclas (29)

*hourly oduct* (30)

1.131367102 1.023607638 -9.52%

\_ P /m <sup>r</sup> \$ <sup>+</sup> <sup>=</sup> *hourly CAPEX hourly OPEX*

calculated as follows:

238 Desalination Updates

NF

Similarly, the hourly OPEX is calculated as follows:

Then the unit product cost of the desalted water is calculated as follows:

<sup>3</sup> Unit product cost, \_ \_

product cost of NF-MSF-DM is 9.5% lower than that of the conventional MSF-BR.

**Levelization cost, \$/m3 MSF-BR NF-MSF-DM % diff** Interest rate 0.07 0.07 - Life span 20 20 - Amortization factor 0.094392926 0.094392926 - Annual investment 192468.7044 215680.4949 12.06% Hourly production 208.07 208 -0.03% Hourly investment 24.41257032 27.35673451 12.06% Specific CAPEX 0.117328641 0.131522762 12.10% Specific OPEX 1.014038462 0.613605769 -39.49% Total 1.131367102 0.745128531 -34.14%

Specific OPEX, NF - 0.062023194 - Specific CAPEX, NF - 0.08493315 - Specific CAPEX, NF-MSF-DM 0.117328641 0.216455912 84.49% Specific CAPEX, NF-MSF-DM 1.014038462 0.807151725 -20.40%

**Table 10.** Levelized cost of MSF and NF-MSF-DM configurations [14]

The levelized cost of capital purchased components and operating invested (chemicals, steam, electricity, O&M) to produce water is calculated as shown in Table (10). The specific OPEX of the MSF-DM is 34% lower than that of the conventional MSF-BR. The specific CAPEX of the MSF-DM is 12% higher than that of the conventional MAF-BR. However, the sum of the total cost invested using the MSF-DM is 34% lower than that of the conventional MSF-BR. Due to adding of the NF system, the specific OPEX of NF-MSF-DM is 20% lower that of the conven‐ tional one, while the specific CAPEX increased by 84% as shown Table (10). The total unit

In this section, technical approach to consider the impact of TBT and varying FO recovery on the process performance is presented. The VDS software [18] will be used as a powerful simulation tool. In this program, a reference MSF plant of 16.2 MIGD working at TBT=111°C is simulated. The performance ratio, distillate production, concentration and flow rates, and temperatures of all streams are calculated. The software is adapted and developed to consider the FO membrane. The hybrid MSF-FO is simulated at fixed brine recycle flow rate and brine concentration (draw solution) and by varying the FO recovery ratio with the TBT. For a fixed performance ratio, the distillate of MSF (D) and the required heat transfer surface area (A) are calculated at different operating conditions. For comparison, the specific heat transfer area (SA) is calculated as:

$$\text{SA} = \frac{A}{D} \text{m}^2/\text{MIGID} \tag{31}$$

For the same seawater and draw solution flow rate across the FO membrane, the permeate (Dm) and the membrane area (Am) are calculated at different recovery ratios. For comparison, the specific area of FO membrane (SA) is calculated as follows:

$$\text{SA} = \frac{A\_m}{D\_m} \text{m}^2/\text{MIGD} \tag{32}$$

At certain FO recovery ratio, the reduction in the Ca+ ions in the MSF feed is calculated and compared to the reference MSF process, which was operated without the FO process. The potential of CaSO4 scale formation in the MSF feed after dilution is estimated at different TBTs (115-135°C) using the Skillman index.

Using VDS, all process stream characteristics are determined (mass, temperature, pressure, entropy, and rated cost), and the heat transfer surface area (number of tubes), evaporator size, internal dimensions, and pumps are sized. So, a detailed CAPEX analysis is performed and estimated. The VDS software calculates the heating steam consumption rate and the consumed chemicals (anti scales, anti-foam, and chlorination), as well as the pumping power (OPEX items). The price of electricity and heating steam is estimated and calculated as illustrated in [18]. Then the final tariff of water unit cost is obtained.

Figure 15 shows the reduction in Ca+ ions in the feed of MSF desalination system at different recovery ratios of the FO membrane system. The reduction of Ca ions increases as the FO recovery ratio increases. At 40% recovery ratio, the reduction in Ca ions is calculated as 20%.

Figure 16 shows that the FO membrane water flux decreases as a result of the increase in the FO recovery ratio. The membrane flux decreases at higher recovery ratio due to dilution of the draw solution, which decreases the osmotic pressure driving force.

Figure 17 shows that the specific membrane area increases as the recovery ratio increases; this is due to lower water flux per unit area at higher recovery ratio. The higher value of specific membrane area is reflected in higher capital cost.

Figure 15: % reduction in Ca+ ions in MSF feed after FO dilution.

**% recovery ratio of FO membrane system**

**Figure 15.** reduction in Ca+ ions in MSF feed after FO dilution Figure 15: % reduction in Ca+ ions in MSF feed after FO dilution. 0 10 15 20 25 30 35 40 45

Figure 16: FO membrane flux at FO membrane recovery ratio variation.

**% recovery ratio of FO membrane**

 80,000 **Specific membrane area Figure 16.** FO membrane flux at FO membrane recovery ratio variation Figure 16: FO membrane flux at FO membrane recovery ratio variation.

Figure 17: Membrane area variation with FO recovery ratio variation.

Figure 17: Membrane area variation with FO recovery ratio variation.

**Figure 17.** Membrane area variation with FO recovery ratio variation

Figure 18 shows the Skillman index (SI) at different TBTs and the variation of FO recovery ratio. The SI of calcium sulphate solubility in case of the reference MSF plant without FO operating at TBT=111°C is calculated as 1.33. As the calculated SI of traditional MSF is greater than 1, the precipitate of calcium sulphate occurs. However, in practical MSF plant anti-scalant is used to disperse the crystalized scale. The SI=1.33 is considered as the reference for com‐ parison; the value above 1.33 indicates scale formation, while the lower value indicates safe operation. As shown in Figure 5, the SI at different TBTs decreases as the FO recovery ratio increases. This is due to the increase in the removal of divalent ions. Figure 5 shows that the Skillman index increases as TBT increases. The MSF can operate at TBT=135 safely without scale problems at an FO recovery ratio of 40%. MSF at TBT=130°C can operate safely at a recovery ratio of 35%. Also, the MSF at TBT=125°C can operate safely at a recovery ratio of 30%. The MSF can operate safely at TBT=120°C and FO recovery ratio of 25%, while the MSF can operate safely at TBT=115°C and FO recovery ratio of 20%.

For the same performance ratio of MSF (PR=9), the reduction in the specific heat transfer surface area of MSF is calculated at different TBTs and different FO recovery ratios as shown in Figure 6. This figure shows that the reduction in specific heat transfer of MSF increases as the TBT increases. The increase of TBT resulting in increase of the logarithmic mean temperature difference between hot and cold streams of MSF, in turn, reduces the heat transfer area. Figure 6 shows that the reduction in SA slightly increases with the increase of the FO recovery ratio.

Figure 18: Skillman index at different MSF TBT and different FO recovery ratio.

**Figure 18.** Skillman index at different MSF TBTs and different FO recovery ratios Figure 18: Skillman index at different MSF TBT and different FO recovery ratio.

Figure 15: % reduction in Ca+ ions in MSF feed after FO dilution.

**Jw, LMH**

**Jw, LMH**

Figure 15: % reduction in Ca+ ions in MSF feed after FO dilution.

**Jw, LMH**

ions in MSF feed after FO dilution Figure 15: % reduction in Ca+ ions in MSF feed after FO dilution.

**% reduction of Ca, ions** 

**% reduction of Ca, ions** 

**% reduction of Ca, ions** 

10 15 20 25 30 35 40 45

**% recovery ratio of FO membrane system**

10 15 20 25 30 35 40 45

**% recovery ratio of FO membrane system**

10 15 20 25 30 35 40 45

**% recovery ratio of FO membrane system**

**the MSF feed**

0.0

10.0

5.0

0.0

**LMH**

**Figure 16.** FO membrane flux at FO membrane recovery ratio variation

0.0

5.0



 - 20,000 40,000 60,000 80,000

**Figure 17.** Membrane area variation with FO recovery ratio variation

 20,000 40,000 60,000 80,000

20,000

40,000

60,000

80,000

**m2/MIGD**

**m2/MIGD**

**m2/MIGD**

5.0

15.0

10.0

10.0

15.0

**Membrane water flux,** 

**Membrane water flux,** 

**Figure 15.** reduction in Ca+

240 Desalination Updates

**LMH**

**LMH**

**Membrane water flux,** 

15.0

**% reduction in Ca ions in** 

**% reduction in Ca ions in** 

**the MSF feed**

**the MSF feed**

**% reduction in Ca ions in** 

Figure 16: FO membrane flux at FO membrane recovery ratio variation.

Figure 16: FO membrane flux at FO membrane recovery ratio variation.

Figure 16: FO membrane flux at FO membrane recovery ratio variation.

**Specific membrane area**

**Specific membrane area**

**Specific membrane area**

0 20 40 60

**% recovery ratio of FO membrane**

0 20 40 60

**% recovery ratio of FO membrane**

0 20 40 60

**% recovery ratio of FO membrane**

Specific…

Specific…

Specific…

Figure 17: Membrane area variation with FO recovery ratio variation.

Figure 17: Membrane area variation with FO recovery ratio variation.

Figure 17: Membrane area variation with FO recovery ratio variation.

Figure 18 shows the Skillman index (SI) at different TBTs and the variation of FO recovery ratio. The SI of calcium sulphate solubility in case of the reference MSF plant without FO operating at TBT=111°C is calculated as 1.33. As the calculated SI of traditional MSF is greater than 1, the precipitate of calcium sulphate occurs. However, in practical MSF plant anti-scalant is used to disperse the crystalized scale. The SI=1.33 is considered as the reference for com‐ parison; the value above 1.33 indicates scale formation, while the lower value indicates safe

0 10 20 30 40 50

**% recovery ratio of FO membrane**

0 10 20 30 40 50

**% recovery ratio of FO membrane**

0 10 20 30 40 50

**% recovery ratio of FO membrane**

**Figure 19.** Reduction in the heat transfer area (CAPEX) at different FO recover ratios and TBTs

For the existing MSF plant, at TBT=130°C, the production will increase up to 30% as shown in Figure 19. The question is whether or not the existing material can withstand the 135°C temperature.

Figure (20): Basic flow diagram of CSP-NF-MSF-DBM pilot plant [13].

Figure (20): Basic flow diagram of CSP-NF-MSF-DBM pilot plant [13].

Figure 19: Reduction in the heat transfer area (CAPEX) at different FO recover ratio & TBTs.

Figure 19: Reduction in the heat transfer area (CAPEX) at different FO recover ratio & TBTs.

Figures 17, 18, and 19 indicate that it is beneficial to work at higher TBT to reduce the heat transfer area or to increase the production of the existing plant of MSF (CAPEX/OPEX reduction); however, this requires higher recovery ratio of the FO membrane system, which requires higher membrane area (CAPEX increase). So, an economical evaluation and compro‐ mise to reach the trade-off point is still required.

The existing capacity of water desalination plants in Qatar is approximately 1.5 Mm3 /day. MSF represent the main technology in Qatar. The make-up of seawater feed is chemically treated (anti-scalant, anti-foulant, and sodium sulphate) before being introduced to the heat recovery section. The amount of make-up flow rate is 3 times of the water production capacity, which is 4.5 Mm3 /day. As shown in Table (11), the chemical cost is 2.35 M\$/year.


**Table 11.** Chemical cost analysis of thermal desalination plant

Thus, it can be concluded that the integrated FO as a pretreatment unit to the seawater feed to the existing MSF desalination plant in Qatar is technically visible in terms of production capacity increase and chemical consumption decrease. However, cost analysis is required to balance the OPEX reduction with the addition caped of the FO membrane unit. Integrating FO to the existing MSF and using the 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 at TBT=130°C, the production of the existing MSF plant increases by 20%. The OPEX analysis showed that an amount of 2.3 M\$/year of chemical cost can be saved if the FO is deployed to the existing MSF plant in Qatar. The tradeoff point between the additional CAPEX of the FO membrane system and the savings in OPEX will be considered under different operating condition in the present work.

### **5. Experimental study of hybrid NF-MSF**

The process design and simulation for the test pilot is developed to prepare specifications of different components. Some units are manufactured by an Egyptian contractor, while others are purchased from vendors. The site is prepared where civil work and foundation are constructed. The test pilot components are installed and assembled, and finally, individual commissioning for each component is carried out. The site is located at the "Wadi El-Natroun" remote area, which is almost 150 km south-west of Alexandria city (Egypt). The site belongs to Alexandria University.

TBT=135 TBT=130 TBT=125 TBT=120 TBT=115

Without FO, TBT=111 TBT=135 TBT=130 TBT=125 TBT=120

Figure 18: Skillman index at different MSF TBT and different FO recovery ratio.

0 5 10 15 20 25 30 35 40 45 50

**FO recovery ratio**

Figure 19: Reduction in the heat transfer area (CAPEX) at different FO recover ratio & TBTs.

10 15 20 25 30 35 40 45 50

**FO recovery ratio**

Figure (20): Basic flow diagram of CSP-NF-MSF-DBM pilot plant [13]. **Figure 20.** Basic flow diagram of CSP-NF-MSF-DBM pilot plant [13]

0.9 1 1.1 1.2 1.3 1.4 1.5 1.6 1.7

> 0.00 0.10 0.20 0.30 0.40

**Reduction in specific** 

**heat transfer area**

**Skillman index**

Figure 20 shows the pilot test of the solar energy and desalination units. The concentrated solar parabolic trough with thermal energy storage facility provides the necessary heating to generate the required steam of the MSF desalination unit. The system is also equipped with a backup boiler for steam compensation. Solar PV and wind turbine (not present in Figure 20) are installed and run separately. However, in this phase, diesel engine is used to provide the pump electricity until the match and synchronization between the PV and the wind turbine are finalized.

#### **5.1. MSF desalination unit**

Figures 17, 18, and 19 indicate that it is beneficial to work at higher TBT to reduce the heat transfer area or to increase the production of the existing plant of MSF (CAPEX/OPEX reduction); however, this requires higher recovery ratio of the FO membrane system, which requires higher membrane area (CAPEX increase). So, an economical evaluation and compro‐

represent the main technology in Qatar. The make-up of seawater feed is chemically treated (anti-scalant, anti-foulant, and sodium sulphate) before being introduced to the heat recovery section. The amount of make-up flow rate is 3 times of the water production capacity, which

/day. MSF

The existing capacity of water desalination plants in Qatar is approximately 1.5 Mm3

/day. As shown in Table (11), the chemical cost is 2.35 M\$/year.

**Chemical Dosing rate, ppm Consumption, kg/day \$/year** Anti-scalant 2.5 2.81E+03 2.28E+06 Anti-foam 0.1 1.35E+01 3.08E+04 Sodium sulfate 0.5 2.25E+02 4.11E+04 Total 2.35E+06

Thus, it can be concluded that the integrated FO as a pretreatment unit to the seawater feed to the existing MSF desalination plant in Qatar is technically visible in terms of production capacity increase and chemical consumption decrease. However, cost analysis is required to balance the OPEX reduction with the addition caped of the FO membrane unit. Integrating FO to the existing MSF and using the 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 at TBT=130°C, the production of the existing MSF plant increases by 20%. The OPEX analysis showed that an amount of 2.3 M\$/year of chemical cost can be saved if the FO is deployed to the existing MSF plant in Qatar. The tradeoff point between the additional CAPEX of the FO membrane system and the savings in OPEX

The process design and simulation for the test pilot is developed to prepare specifications of different components. Some units are manufactured by an Egyptian contractor, while others are purchased from vendors. The site is prepared where civil work and foundation are constructed. The test pilot components are installed and assembled, and finally, individual commissioning for each component is carried out. The site is located at the "Wadi El-Natroun" remote area, which is almost 150 km south-west of Alexandria city (Egypt). The site belongs

will be considered under different operating condition in the present work.

mise to reach the trade-off point is still required.

**Table 11.** Chemical cost analysis of thermal desalination plant

**5. Experimental study of hybrid NF-MSF**

to Alexandria University.

is 4.5 Mm3

242 Desalination Updates

The MSF pilot unit consists of 28 flashing chambers with 28 connected condensers as shown in Figures 21.a and 21.b. The stages are arranged in double deck to reduce the foot print. There are four sets; each set consists of 7 stages. MSF chambers are equipped with glass windows for monitoring of the flashing process. The shell material of MSF is fabricated from 2 mm-thick stainless steel 316L. The flash chamber is 0.5 m in length and 0.5 m in width, while the height is 1.0 m. The condenser tube is 8 mm in diameter, 6 m long, and made of stainless steel (0.7 mm thick). The number of tubes is 2 per condenser, which are arranged in multi-pass inside a 0.5 m shell length. The unit is manufactured in Egypt and assembled at the project site.

The orifice opening area is controlled using gate valve, which is located between the flash chambers. The inter-stage valves controls the inter-stage flow rates to guarantee the brine flashing at each stage. The splash plate is designed just above of the inlet opening to reduce the carry-over. A demister is placed near the vapor outlet vapor pipes to reject the brine carryover before going to the condenser. The shells are insulated to minimize energy losses. In addition to the brine heater, different supporting systems are added including vacuum system and chemical injection systems. The vacuum system has control valves at each stage to adjust the venting rate of non-condensable gases (NCG) and the stage pressure.

The MSF is the main subsystem where distillation is produced using the flashing process. Different instrumentations are installed to measure and record the temperatures, pressures, and flow rates as shown in Figure 20. In the heating section, steam input and output temper‐ atures, in addition to pressure and flow, are measured using proper transducers. All chambers are equipped with temperature and pressure indicators. The first and last chambers are equipped with temperature transmitter (TT) and pressure transmitter (PT), and the two

a- Double deck MSF desalination test pilot unit b- Process flow diagram of Double deck MSF

a- VDS simulation results of CSP[13] b- Solar and absorbed energy [13]

**Figure 21.** MSF desalination unit with double deck [13]

additional movable PT and TT are supplied to be inserted in the chambers of the operator choice. Input seawater flow and output brine and distilled water flow rates are measured using flow transmitters.

#### **5.2. Concentrated solar power (CSP) system**

Four modules of solar concentrator (parabolic trough) are purchased and assembled in series at the site of the project as shown in Figure 22. Each module is 3.6 m in length and 1.524 in width. The collector area per module is 5.6 m2, while the collector reflective area is 5 m2 . The assembled collector length becomes 14.6 m, while the total area is 22.4 m2 . The receiver absorptivity is 0.92, the mirror reflectivity is 0.91, while the receiver emittance is 0.23. The blackcoated pipes are 1.0 inch in diameter and placed in 2.0-inch-diameter glass pipes to minimize convection losses. The concentrators have a tracking system and were placed east-west and facing south.

The CSP system contains a steam generator to supply the MSF brine heater with the required heating steam. Thermal oil is circulated through the collecting pipes, gains the solar thermal energy, and flows through the steam generator and energy storage tank. The steam generator

**Figure 22.** Four modules of the concentrated solar collector in series [13]

consists of shell and tube and has a separate vapor header. The shell diameter is 10 inches and is 2 m in length. The hot oil passes through tubes, while the water flows through the shell. The tube length is 4 m, and the diameter is 6 mm; the number of tubes is 24, which are arranged in two passes. The CSP system is instrumented with temperature transmitters (TT), flow meters (FT), and pressure transmitters (PT), as shown in Figure 20, to monitor the temperatures, flow rates, and pressure in both steam and oil loops.

#### **5.3. NF pretreatment**

. The

. The receiver

additional movable PT and TT are supplied to be inserted in the chambers of the operator choice. Input seawater flow and output brine and distilled water flow rates are measured using

a- VDS simulation results of CSP[13] b- Solar and absorbed energy [13]

a- Double deck MSF desalination test pilot unit b- Process flow diagram of Double deck MSF

**Solar Energy, W/m2**

Input…

**Day time**

Four modules of solar concentrator (parabolic trough) are purchased and assembled in series at the site of the project as shown in Figure 22. Each module is 3.6 m in length and 1.524 in width. The collector area per module is 5.6 m2, while the collector reflective area is 5 m2

absorptivity is 0.92, the mirror reflectivity is 0.91, while the receiver emittance is 0.23. The blackcoated pipes are 1.0 inch in diameter and placed in 2.0-inch-diameter glass pipes to minimize convection losses. The concentrators have a tracking system and were placed east-west and

The CSP system contains a steam generator to supply the MSF brine heater with the required heating steam. Thermal oil is circulated through the collecting pipes, gains the solar thermal energy, and flows through the steam generator and energy storage tank. The steam generator

assembled collector length becomes 14.6 m, while the total area is 22.4 m2

flow transmitters.

244 Desalination Updates

facing south.

**5.2. Concentrated solar power (CSP) system**

**Figure 21.** MSF desalination unit with double deck [13]

Figure 23.a shows the P&I diagram of the NF system. The system consists of dual media filter, cartridge filter, high-pressure pumps, chemical injection pumps, and nano-filtration (NF) membrane. One dual media filter vessel is installed with a specified feed flow rate of 1.5 ton/hr and 3.5 ton/hr for back wash. The vessel contains sand, gravel, and anthracite. The cartridge filter of 5 micron is installed after high pressure pump and just before the membrane section. The membrane section consists of 4 pressure vessels running in parallel; each vessel contains one membrane element of NF270 4040 type. The whole NF system, except feed, permeate, and brine tanks, is placed inside one container with its control panel, as shown in Figure 23.b. For water salinity, samples are collected periodically to measure the conductivity using a mobile conductivity meter.

a- Process Flow Diagram of NF system b- Photo of NF room

Figure 23: Nano Filtration (NF) system [13]. **Figure 23.** Nano-filtration (NF) system [13]

#### **5.4. NF test performance**

The NF system testing is carried out using the site brackish water (TDS=2000 ppm). A math‐ ematical model of the NF membrane is developed and verified against typical operating NF unit data using the VDS software developed by the authors [13-15]. The VDS simulation results of the NF system were derived at different feed pressures of 8 and 10 bars and compared with experimental results as shown in Table (12). The NF performance was carried out and assessed by the recovery ratio and salt rejection. The recovery ratio (permeate/feed) increases as the feed pressure increases. The salt rejection (1- (permeate salinity/feed salinity)) is calculated as shown in Table (12). The salt rejection decreases as the feed pressure increases due to the increase in permeate salinity. The measured recovery ratio is slightly lower than the simulation results, although the salt rejection determined in the experiment is lower than that of the simulation. The differences between the measured values of permeate flow, salinity, recovery ratio, salt rejection, and simulation results are within the acceptable range.


**Table 12.** Typical NF experimental results compared with the VDS results [13]

#### **5.5. Concentrated solar power (CSP) test performance**

The CPS system, including the solar collector and steam generator, is simulated using the VDS program. The mass and heat balance equations of the solar collector, steam generator, and pumps are developed. The oil and water thermo-physical property correlations at different temperatures are considered in the program. The characteristic surface of collector reflectivity, receiver emission and absorptivity, and glass tube material transmittance are specified in the VDS program. The specifications of the solar collector and steam generator are defined and fed to the program. The measured weather conditions (solar intensity, ambient temperature, and wind velocity) at each hour are fed to the program. The duration time starts at 7:00 AM and ends at 8:00 PM with 1 hour step. a- Double deck MSF desalination test pilot unit b- Process flow diagram of Double deck MSF

#### **Figure 24.** ADD CAPTION

a- Process Flow Diagram of NF system b- Photo of NF room

The NF system testing is carried out using the site brackish water (TDS=2000 ppm). A math‐ ematical model of the NF membrane is developed and verified against typical operating NF unit data using the VDS software developed by the authors [13-15]. The VDS simulation results of the NF system were derived at different feed pressures of 8 and 10 bars and compared with experimental results as shown in Table (12). The NF performance was carried out and assessed by the recovery ratio and salt rejection. The recovery ratio (permeate/feed) increases as the feed pressure increases. The salt rejection (1- (permeate salinity/feed salinity)) is calculated as shown in Table (12). The salt rejection decreases as the feed pressure increases due to the increase in permeate salinity. The measured recovery ratio is slightly lower than the simulation results, although the salt rejection determined in the experiment is lower than that of the simulation. The differences between the measured values of permeate flow, salinity, recovery

**VDS Exp % diff VDS simulation Exp % diff**

Figure 23: Nano Filtration (NF) system [13]. **Figure 23.** Nano-filtration (NF) system [13]

ratio, salt rejection, and simulation results are within the acceptable range.

Permeate flow rate, ton/hr 0.67 0.6 -10% 0.85 0.76 -11% Permeate salinity, ppm 648 600 -7% 761 650 -15% Brine flow rate, ton/hr 0.65 0.725 12% 0.52 0.6 15% Brine salinity, ppm 3395 3158 -7% 4027 3760 -7% Recovery ratio, % 51 45 -12% 62 55 -11% Salt rejection, % 67.6 70 4% 62 68 10%

Feed pressure, bar 8 10 Feed flow rate, ton/hr 1.325 1.375 Feed salinity, ppm 2000 2000

**Table 12.** Typical NF experimental results compared with the VDS results [13]

**5.4. NF test performance**

246 Desalination Updates

Figure 24.a shows the interface of the VDS program results at 4:00 PM. The oil mass flow rate and the temperature at inlet and outlet for both the solar collector and steam generator are presented. The collected energy is transferred to the steam generator to generate 3.8 kg/hr of saturated steam at 113.8°C. The solar intensity (I) and calculated absorbed energy by the receiver are shown in Figure 24.b. The difference is noticeable at mid-day time. Figure 24.a: VDS simulation results of CSP[13] 24.b- Solar and absorbed energy [13]

a- temperature rise in CSP b- Temperature drop in steam generator

**Day time**

Figure 25: Oil temperature variation through CSP trough & Steam generator [13]. **Figure 25.** Oil temperature variation through CSP trough and steam generator [13]

Figure 25.a shows comparison between the VDS simulation and experimental results of oil temperature rise through the solar collector during day time. The oil temperature difference increases as the solar intensity increases, while the maximum difference at mid-day reached 25°C. Figure 25.b shows a comparison between the simulation and experimental results of the oil temperature drop in the steam generator unit. The maximum heat transfer occurs during mid-day, and the maximum temperature drop is 14°C. It is similarly noticed that at day time, the temperature drop in steam generator unit is less than the temperature rise in the solar collector. This means that part of the gained energy in the collector is absorbed in the steam generator, and the remaining is maintained with the outlet oil stream from the steam generator and comes back the collector. This explains the increase of oil temperature at the concentrated solar collector inlet at day time.

The CSP system average efficiency () is calculated as the average useful gained power/average solar input power:

$$\eta\_{\rm CSP} = \frac{\dot{m}C\_p \Big|\_{ol} (T\_{o,all} - T\_{t,all})}{I \times A\_{\rm CSP}} \tag{33}$$

Figure 26: Solar Collector efficiency [13] Figure 27: Generated steam temperature [13]

**Temperature, C**

**Temperature, C**

Sim… Exp.

**Day time, hr**

**TBT**

**Day time, hr**

Simu… Exp.

Figure 28: Generated steam flow rate, [13] Figure 29: TBT variation, [13]

Simulati on

Figure 30: MSF distillate productivity with TBT variation [13].

**TBT, C**

y = 0.0813x1.3917 y = 0.1195x R² = 0.68291.27 R² = 0.5919

**MSF distillate**

Figure 26 shows the simulation and experimental results of the collector efficiency at day time. The collector efficiency decreases during day time due to the increase in the average oil temperature, which increases the energy loss to the ambient. The experimental collector efficiency shows relatively low value than that of the simulation collector efficiency due to: i) inaccurate tracking system that could not follow the sun movements accurately, and ii) the inefficient concentrated tube location in the CSP focus and possible convection loss.

Figure 27 shows comparison between the simulated and the measured generated steam temperatures. The water inlet and steam exit valve remain closed while the oil valves are open to allow energy transfer from oil to heat the enclosed water in the boiler. The water feed and steam valves are opened when the water temperature reaches 77°C. The generated steam temperature increases as the solar intensity increases, and the maximum temperature reached is 110°C at mid-day time.

**Generated steam flow rate**

**Day time, hr**

**Distillate, kg/hr**

**Steam flow rate, kg/hr**

Simulati on

**Figure 26.** ADD CAPTION

**Temperature, C**

**Temperature, C**

y = 0.1195x R² = 0.68291.27 R² = 0.5919

Sim… Exp.

**Day time, hr**

**TBT**

**Day time, hr**

Simu… Exp.

Figure 26: Solar Collector efficiency [13] Figure 27: Generated steam temperature [13]

**Figure 27.** ADD CAPTION

**Day time**

**Generated steam flow rate**

**CSP Efficiency**

**Day time, hr**

**Day time**

**Generated steam flow rate**

**CSP Efficiency**

Exprime ntal

Simulati on

Exprime ntal

Simulati on

**CSP Efficiency**

**Steam flow rate, kg/hr**

**CSP Efficiency**

**Steam flow rate, kg/hr**

Figure 26: Solar Collector efficiency [13] Figure 27: Generated steam temperature [13]

Figure 25.a shows comparison between the VDS simulation and experimental results of oil temperature rise through the solar collector during day time. The oil temperature difference increases as the solar intensity increases, while the maximum difference at mid-day reached 25°C. Figure 25.b shows a comparison between the simulation and experimental results of the oil temperature drop in the steam generator unit. The maximum heat transfer occurs during mid-day, and the maximum temperature drop is 14°C. It is similarly noticed that at day time, the temperature drop in steam generator unit is less than the temperature rise in the solar collector. This means that part of the gained energy in the collector is absorbed in the steam generator, and the remaining is maintained with the outlet oil stream from the steam generator and comes back the collector. This explains the increase of oil temperature at the concentrated

The CSP system average efficiency () is calculated as the average useful gained power/average

, , ( )

*p o oil i oil*

*CSP*

Figure 26 shows the simulation and experimental results of the collector efficiency at day time. The collector efficiency decreases during day time due to the increase in the average oil temperature, which increases the energy loss to the ambient. The experimental collector efficiency shows relatively low value than that of the simulation collector efficiency due to: i) inaccurate tracking system that could not follow the sun movements accurately, and ii) the

Figure 27 shows comparison between the simulated and the measured generated steam temperatures. The water inlet and steam exit valve remain closed while the oil valves are open to allow energy transfer from oil to heat the enclosed water in the boiler. The water feed and steam valves are opened when the water temperature reaches 77°C. The generated steam temperature increases as the solar intensity increases, and the maximum temperature reached

**CSP Efficiency**

Exprime ntal

Simulati on

**Day time**

**Generated steam flow rate**

**Day time, hr**

**Distillate, kg/hr**

Figure 26: Solar Collector efficiency [13] Figure 27: Generated steam temperature [13]

**Temperature, C**

**Temperature, C**

Sim… Exp.

Figure 28: Generated steam flow rate, [13] Figure 29: TBT variation, [13]

Simulati on

Figure 30: MSF distillate productivity with TBT variation [13].

**TBT, C**

y = 0.0813x1.3917 y = 0.1195x R² = 0.68291.27 R² = 0.5919

**MSF distillate**


(33)

**Day time, hr**

**TBT**

**Day time, hr**

Simu… Exp.

.

<sup>=</sup> ´

*oil*

inefficient concentrated tube location in the CSP focus and possible convection loss.

*mC T T*

*I A*

h

*CSP*

**CSP Efficiency**

**Steam flow rate, kg/hr**

solar collector inlet at day time.

solar input power:

248 Desalination Updates

is 110°C at mid-day time.

**Figure 26.** ADD CAPTION

Figure 28: Generated steam flow rate, [13] Figure 29: TBT variation, [13] **Figure 28.** Generated steam flow rate [13] Figure 26: Solar Collector efficiency [13] Figure 27: Generated steam temperature [13]

Simulati on

Figure 28: Generated steam flow rate, [13] Figure 29: TBT variation, [13]

Figure 30: MSF distillate productivity with TBT variation [13].

**TBT, C**

**Distillate, kg/hr**Figure 28: Generated steam flow rate, [13] Figure 29: TBT variation, [13] **Figure 29.** TBT variation [13]

**Day time, hr**

Figure 30: MSF distillate productivity with TBT variation [13]. **TBT, C** Figure 30: MSF distillate productivity with TBT variation [13]. **TBT, C** y = 0.0813x1.3917 y = 0.1195x R² = 0.68291.27 R² = 0.5919 **Distillate, kg/hr MSF distillate** Simulati on The steam valve is opened at 1:00 PM at a steam flow rate of 4.3 kg/hr. The generated steam is directed to the MSF desalination unit as a heating source. The condensate steam in the brine heater of MSF is fed back to the steam generator. The amount of generated steam flow rate decreased linearly, as shown in Figure 28, due to the decrease in the solar collector efficiency. The measured generated steam flow rate shows lower values than the simulation results due to the thermal losses encountered through insufficient insulation of steam generator and throughout the connection pipe between the CSP outlet and steam generator. As shown in Figure 28, the operation of the steam generation extends up to 11:00 PM due to the heat storage in the CSP system.

Figure 29 shows the simulation and experimental values of TBT variation during day time. As the CSP steam condenses in the brine heater of MSF, the TBT rises up due to the gained energy of the latent heat. Under the ambient and operating condition in June 2012, the midday TBT reaches up to 100°C while the CSP steam condenses at 106°C, that is, 6°C temperature differ‐ ence is maintained. **CSP Efficiency CSP Efficiency** Exprime **Temperature, C**

Simu… Exp.

#### **5.6. New MSF with de-aerator and brine mix (DBM)** ntal

The permeate water of the NF system is directed to the de-aeration and brine mix tower, where the feed is sprayed for oxygen removal. The deaerated water is mixed with parts of the brine blow down, then it is pumped to the MSF condensers. The brine mix feed absorbs latent heat energy in condensers before passing though the brine heater where the brine reaches its top temperature (TBT). The brine is then directed to the first flash chamber where flashing process occurs and vapor releases. The released vapor condenses to form product water. The flashing process occurs in the successive stages untilthe last stage is reached, where the un-flashedbrine exits as brine blown down. The condensate of all stages is collected and directed to the water product tank. The brine level is adjusted above the interconnecting pipes (inter-stage gates) to guarantee the sealing of the flash chambers. Figure 26: Solar Collector efficiency [13] Figure 27: Generated steam temperature [13] **Day time Day time, hr Steam flow rate, kg/hr Generated steam flow rate** Simulati **Temperature, C TBT**

Under the same feed saline water flow rate (NF permeate) of 370 kg/hr and feed temperature of 27°C and controlling the brine mix ratio at 20-70% of the MSF brine blow down, the distil‐ late wateris measured and recorded at different TBTs, as shown in Figure 30, as compared with design calculated values. The pressure of saline water before the first chamberis controlled and fixedat anabsolute value of 1.5 bars (above saturationconditions) bypartially closing the valve. Also,the orifices among chambers are controlled by partially closing the valve between the two successive chambers. The in-tube water velocity is controlled at 2 m/s. Figure 28: Generated steam flow rate, [13] Figure 29: TBT variation, [13] **Day time, hr** on **Day time, hr** Sim… Exp.

**MSF distillate**

Figure 30: MSF distillate productivity with TBT variation [13]. **Figure 30.** MSF distillate productivity with TBT variation [13]

The rates of both the design and the measured distilled water increase as a result of the TBT increases as shown in Figure 30. The amount of distillate is lower than the expected; this may be due to the partial loss of flashed vapor through vacuum system and the irreversibility of the flashing process that occurs within the orifices and weirs.

Figure 31 shows the design and the experimental GOR of the MSF variation with TBT. This is defined as the ratio between the distillate flow rate and the heating steam consumption,. The

Techno-Economics of Hybrid NF/FO with Thermal Desalination Plants http://dx.doi.org/10.5772/60207 251

Simulation

Figure 31: The GOR of MSF system variation with TBT [13]. **Figure 31.** The GOR of MSF system variation with TBT [13]

**GOR**

Figure 29 shows the simulation and experimental values of TBT variation during day time. As the CSP steam condenses in the brine heater of MSF, the TBT rises up due to the gained energy of the latent heat. Under the ambient and operating condition in June 2012, the midday TBT reaches up to 100°C while the CSP steam condenses at 106°C, that is, 6°C temperature differ‐

**Temperature, C**

**Day time, hr**

**TBT**

**Day time, hr**

Simu… Exp.

The permeate water of the NF system is directed to the de-aeration and brine mix tower, where the feed is sprayed for oxygen removal. The deaerated water is mixed with parts of the brine blow down, then it is pumped to the MSF condensers. The brine mix feed absorbs latent heat energy in condensers before passing though the brine heater where the brine reaches its top temperature (TBT). The brine is then directed to the first flash chamber where flashing process occurs and vapor releases. The released vapor condenses to form product water. The flashing process occurs in the successive stages untilthe last stage is reached, where the un-flashedbrine exits as brine blown down. The condensate of all stages is collected and directed to the water product tank. The brine level is adjusted above the interconnecting pipes (inter-stage gates) to

Figure 26: Solar Collector efficiency [13] Figure 27: Generated steam temperature [13]

Under the same feed saline water flow rate (NF permeate) of 370 kg/hr and feed temperature of 27°C and controlling the brine mix ratio at 20-70% of the MSF brine blow down, the distil‐ late wateris measured and recorded at different TBTs, as shown in Figure 30, as compared with design calculated values. The pressure of saline water before the first chamberis controlled and fixedat anabsolute value of 1.5 bars (above saturationconditions) bypartially closing the valve. Also,the orifices among chambers are controlled by partially closing the valve between the two

> y = 0.0813x1.3917 y = 0.1195x R² = 0.68291.27 R² = 0.5919

**MSF distillate**

**Temperature, C**

Sim… Exp.

Figure 30: MSF distillate productivity with TBT variation [13].

The rates of both the design and the measured distilled water increase as a result of the TBT increases as shown in Figure 30. The amount of distillate is lower than the expected; this may be due to the partial loss of flashed vapor through vacuum system and the irreversibility of

Figure 31 shows the design and the experimental GOR of the MSF variation with TBT. This is defined as the ratio between the distillate flow rate and the heating steam consumption,. The

**TBT, C**

Figure 28: Generated steam flow rate, [13] Figure 29: TBT variation, [13]

successive chambers. The in-tube water velocity is controlled at 2 m/s.

Simulati on

**Distillate, kg/hr**

the flashing process that occurs within the orifices and weirs.

**Figure 30.** MSF distillate productivity with TBT variation [13]

ence is maintained.

250 Desalination Updates

**CSP Efficiency**

**Steam flow rate, kg/hr**

**5.6. New MSF with de-aerator and brine mix (DBM)**

**Day time**

**Generated steam flow rate**

**CSP Efficiency**

Exprime ntal

guarantee the sealing of the flash chambers.

Simulati on

**Day time, hr**

average value of the unit design GOR is 17, which is almost twice of the conventional MSF GOR. The average measured GOR is 15 as shown in Figure 31. The small difference between the measured and designed values of GOR is due to the lower distillate productivity under fixed amount of heating steam flow rate. Figure 31: The GOR of MSF system variation with TBT [13]. **TBT, C** EXP. **MSF Specific power consumption**

The MSF specific power consumption (SPC) is defined as the ratio between the pumping power consumption (kW) and the rate of water distillate (m3 /hr),. Simulation EXP.

**MSF Specific power consumption**

Figure 32: Specific Power Consumption (SPC) of MSF with TBT variation [13]. **Figure 32.** Specific power consumption (SPC) of MSF with TBT variation [13]

Figure 32 shows thatthe SPC decreases as theTBTincreasesmainlydue to the increases in water productivity.The experimental SPC is calculatedbasedon the measureddistillate flow rate and the rated power consumption. The experimental SPC is higher, however, than the design value mainly due to lower experimental distillate for the same saline water feed and may be due to the pressure drop in piping and valves that were not considered properly in the design stage. The SPC of the test pilot unit is relatively higher than that of the commercial value of large scale MSF desalination plant due to the very small test pilot unit productivity.

### **6. Conclusion**

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 atthe same site. Product waterfrom the membrane and thermal systems are usually blended to international standards on water quality. One more step ahead this chapter addresses the role of using FO or NF as a pre-treated method to the existing thermal desalina‐ tion plants. The target of this hybridization is to reduce divalent ions that cause hard-scale deposition at elevated temperatures. The separation of divalent ion enables increasing the desalination process temperature greater than 110°C, which consequently increases the plant performance, increases the productivity, and reduces chemical consumption.

Integrating the NF system with new (MSF-DM) configuration at TBT=130°C, the gain output ratio could be as high as 16, which is double of that for the conventional MSF-BR. The new NF-MSFDM configuration significantly reduces the unit's input thermal energy to suit the use of (the relatively expensive) solar energy as a desalination plant driver.

Simulation results showed that integrating FO to the existing MSF and using the 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 at TBT=130°C, the production of the existing MSF plant increases by 20%. The OPEX analysis showed that an amount of 2.3 M\$/year of chemical cost can be saved if the FO is deployed to the existing MSF plant in Qatar.

Thedesalinationpilottestisbuilttoevaluate theperformance ofthenovelde-aerationandbrine mix (MSF-DBM) configuration at high TBT using the NF membrane. The capacity of the desalination pilot plant is 1.0 m3 /day of water. Comparison between the simulation and the experimental results of the pilot unit subsystems is relatively satisfactory. The newly devel‐ oped NF-MSF-DBM (de-aerator and brine mix) configuration is tested at TBT=100°C, and the GOR is calculated as 15, which is almost twice of the traditional MSF under the same operat‐ ing conditions. Using the new high-performance NF-MSF-DBM and the unit's input thermal energy, which make the integration with (the relatively expensive) RE as a desalination plant driver, is a viable option.

### **Author details**

Abdel Nasser Mabrouk1,2\*, Hassan Fath3 , Mohamed Darwish1 and Hassan Abdulrahim1

\*Address all correspondence to: aaboukhlewa@qf.org.qa abdelnaser.mabrouk@suezu‐ niv.edu.eg

1 Qatar Environmental & Energy Research Institute, Qatar

2 Suez University, Egypt

3 American University, Sharjah, UAE

### **References**

**6. Conclusion**

252 Desalination Updates

the existing MSF plant in Qatar.

desalination pilot plant is 1.0 m3

Abdel Nasser Mabrouk1,2\*, Hassan Fath3

3 American University, Sharjah, UAE

1 Qatar Environmental & Energy Research Institute, Qatar

driver, is a viable option.

2 Suez University, Egypt

**Author details**

niv.edu.eg

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 atthe same site. Product waterfrom the membrane and thermal systems are usually blended to international standards on water quality. One more step ahead this chapter addresses the role of using FO or NF as a pre-treated method to the existing thermal desalina‐ tion plants. The target of this hybridization is to reduce divalent ions that cause hard-scale deposition at elevated temperatures. The separation of divalent ion enables increasing the desalination process temperature greater than 110°C, which consequently increases the plant

Integrating the NF system with new (MSF-DM) configuration at TBT=130°C, the gain output ratio could be as high as 16, which is double of that for the conventional MSF-BR. The new NF-MSFDM configuration significantly reduces the unit's input thermal energy to suit the use of

Simulation results showed that integrating FO to the existing MSF and using the brine of the

by 20%, which enables increasing the TBT up to 130°C safely. The simulation results show that at TBT=130°C, the production of the existing MSF plant increases by 20%. The OPEX analysis showed that an amount of 2.3 M\$/year of chemical cost can be saved if the FO is deployed to

Thedesalinationpilottestisbuilttoevaluate theperformance ofthenovelde-aerationandbrine mix (MSF-DBM) configuration at high TBT using the NF membrane. The capacity of the

experimental results of the pilot unit subsystems is relatively satisfactory. The newly devel‐ oped NF-MSF-DBM (de-aerator and brine mix) configuration is tested at TBT=100°C, and the GOR is calculated as 15, which is almost twice of the traditional MSF under the same operat‐ ing conditions. Using the new high-performance NF-MSF-DBM and the unit's input thermal energy, which make the integration with (the relatively expensive) RE as a desalination plant

, Mohamed Darwish1

\*Address all correspondence to: aaboukhlewa@qf.org.qa abdelnaser.mabrouk@suezu‐

/day of water. Comparison between the simulation and the

ions in the seawater feed

and Hassan Abdulrahim1

performance, increases the productivity, and reduces chemical consumption.

(the relatively expensive) solar energy as a desalination plant driver.

last stage as a draw solution at a recovery ratio of 35% reduce the Ca+


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