Desalination 307 (2012) 76–86

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Desalination

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Enhancement of existing MSF plant productivity through design modification andchange of operating conditions

A.M. Helal ⁎, A. Al-Jafri, A. Al-YafeaiNational Energy and Water Research Center (NEWRC), Abu Dhabi Water and Electricity Authority (ADWEA), P.O. Box 54111, Abu Dhabi, United Arab Emirates

H I G H L I G H T S

► Capacity enhancement for a conventional and a modified design MSF plant is given.► An increase of 49% can be achieved by increasing the recycle rate and TBT.► Increase in product capacity due to design modification only does not exceed 2%.► Additional pumping capacity for some process streams will be necessary.

⁎ Corresponding author. Tel.: +971 50 5627810; fax:E-mail address: Amhelal@adwea.ae (A.M. Helal).

0011-9164/$ – see front matter © 2012 Elsevier B.V. Allhttp://dx.doi.org/10.1016/j.desal.2012.08.027

a b s t r a c t

a r t i c l e i n f o

Article history:Received 16 April 2012Received in revised form 22 August 2012Accepted 22 August 2012Available online 29 September 2012

Keywords:Multi-Stage FlashCapacity enhancementThermal desalinationDesign modifications

This study investigates the possibility of capacity enhancement of two brine recycle-MSF (Multi-Stage Flash)plants. The first has a conventional design whereas in the other the distillate from the heat recovery section isdiverted, after cooling, to the town water tank. Thus the distillate from the recovery section is prevented fromsuccessive re-flashing in the heat rejection section. Also the distillate corridor from the heat rejection sectionis removed and condensate from heat rejection stages is collected individually into a common header afterflashing and cooling to form a second product stream. Both plants incorporate a nanofiltration unit for thepartial removal of bivalent scale forming ions from the makeup stream to enable operation at elevatedTBT. Under the same operating conditions, it was found that the modified design exceeds the conventionalone only by about 2% increase in distillate capacity. At the maximum recycle rate, 85% of the maximumrecycle pump capacity, the increase in top brine temperature from 110 °C to 130 °C results in productcapacity increase of 49.1 and 49.74% for the conventional and modified plants respectively. In many instancesit was necessary to install additional pumping capacity to one or more of the process streams.

© 2012 Elsevier B.V. All rights reserved.

1. Introduction

The Multi-Stage Flash-MSF desalination process is the dominanttechnology for seawater desalination in the Gulf region. The processis capital and energy intensive and it is always desirable tomaximize plant capacity to reduce the specific capital cost. Increasingexisting plant productivity can be realized through one or more of thefollowing passes:

a- Increasing the top brine temperature, TBT which means an increasein the flashing range.

b- Increasing the recycle stream flow rate. This option overloads therecycle pump and reduces its energy efficiency. In addition, if theTBT is to be increased simultaneously, steam amount to the brineheater as well as steam temperature will have to be increasedaccordingly. In this contest, a novel idea in the process pretreatmentis to couple the MSF process with a nanofiltration membrane unit,

+971 2 6948994.

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NF, for partial elimination of the bivalent scale forming ions, namely:SO4

−−, Ca++ and Mg++as well as HCO3− from part of the makeup

stream. This pretreatment step is expected to enable plant operationat a TBT beyond 125 °C [1]. This elevation in TBT increases the flash-ing range and hence plant productivity without the threat of scaleformation at the high temperature end of the plant. It is alsoexpected that the process will be run at reduced anti-scalantdoses, which will result in a lower cost of chemicals and operationalcost in general.

Summariva et al. [2] reported that increase in productivity ofabout 45% can be achieved through increased flashing range andrecycle flow rate in conjunction with the design modification of theheat rejection section. They suggested the diversion of the distillatestream leaving the heat recovery section of the plant for possibleuse of its enthalpy in a process that can benefit from low gradeheating energy, e.g., an MED, Multi‐Effect Desalination, unit. Thus,the distillate stream from the recovery section is not allowed to con-tinue re-flashing and cooling in the heat rejection section. This mod-ification in the process is accompanied by the removal of the distillate

Table 1Base case plant data.

Geometric data

Heat recovery sectionNo of recovery stages 15No of tubes per stage 2861Tube length 15.9 mTube outside diameter 0.0318 mTube inside diameter 0.0293 mTotal demister area in Ht. Rec. section 389 m2

Heat rejection sectionNo of rejection stages 3No of tubes per stage 2714Tube length 15.9 mTube outer diameter 0.0318 mTube Inside diameter 0.0293 mTube material :stages 1–3, Cu–Ni 66/30/2/2Tube material :stages 4–15, Cu–Ni 90/10

Stage dimensionsStage width — all stages 15.75 mStage length in heat recovery section 3.61 mStage length in heat rejection section 3.9 m

Brine heater dataTotal no of tubes 2700Tube length 17.33 mTube outside diameter 0.0318 mTube inside diameter 0.0293 m

Orifices: rectangular-adjustable height

77A.M. Helal et al. / Desalination 307 (2012) 76–86

corridor from the rejection section and directing the condensate fromeach one of the rejection stages to flashing tanks for cooling beforecollection in a main header. A second product stream is thus formedwhich is then directed to the town water tank. Removal of thedistillate corridor from the heat rejection section exposes more heattransfer area to condense additional amounts of vapor that wouldbe flashed off the brine pools in the rejection stages. Here, the recyclerate or the top brine temperature, or both, would be increased, whichresults in increasing distillate rate from the heat rejection stagescompared to the conventional design. Obviously, pumping the twoproduct streams from the two plant sections will call for the installa-tion of an additional product pump.

The added water capacity is said to have a lower specific cost percubic meter than that resulting from building a new RO, ReverseOsmosis, plant having the same capacity [1].

In any case, modification of existing MSF plants with their conven-tional design, to increase productivity must be considered in the lightof the maximum capacities of the different pumps within the systemsuch as the recycle pump, the distillate product pump, the blowdownpump and the condensate recycle pump.

Again, modifications will include readjustment of the venting sys-tem, increase of steam rate and temperature and capital investmentfor the engineering modification of the evaporator.

The idea is worth investigating to explore its technical feasibilityand further to study its cost effectiveness. The present work focuseson the assessment of the technical feasibility of the proposed modifi-cations to shed lights on its merits and demerits.

Orifice width 15 mOrifice Coefficient 0.6 m

Heights to demistersHeat recovery stages 2.9 mHeat rejection stages 3.15 mTotal demister area in Ht. Rec. section 389 m2

Temperatures and concentrationsFeed seawater temperature 32 °CRecycle temp to brine heater 100.0 °CTop brine temperature 110.0 °CSteam temperature in the brine heater 120 °CTotal number of stages 18Seawater feed concentration 50,000 ppm

2. Conventional MSF desalination plant description

Fig. 1 represents a general schematic of a conventional brine-recycle MSF desalination plant with the design and operational char-acteristics outlined in Table 1. Table 2 shows the maximum capacitiesof the different pumps amongst other operational constraints. Theplant characteristics are very similar to those built by Italimpianti atUmm Al-Nar East in Abu Dhabi, The United Arab Emirates.

The evaporator unit consists of 18 stages arranged in a one passconfiguration. Each stage consists of a flash chamber and a horizontaltube heat exchanger. Stages 1–15 comprise the heat recovery section

Intake seawaterpump

Blow down pump

Reject coolant

Heat recovery section

Makeup

Heat rejection section

Brine heater

De-aerator

Recycle

Product pump

Condensaterecycle pumps

Steam

Fig. 1. A schematic of conventional MSF desalination plant.

Table 2Operational constraints.

Maximum capacity of recycle pump 15,411 ton/hMaximum capacity of Intake pump 16,775 ton/hMaximum capacity of distillate product pump 1600 ton/hMaximum capacity of blow down pump 5555 ton/hMaximum capacity of condensate pump 220 ton/hMaximum velocity in tubes 2.2 m/sMinimum velocity in tubes 1 m/sMaximum vapor velocity in the last rejection stage 8.9 m/sMaximum brine loading 1200 ton/h m widthMaximum blow down concentration 80,000 ppmMaximum TBT, top brine temperature 130 °C

Distillate corridor conditionsMaximum dist rate 0.5 m3/sHeight of distillate orifice 0.20 mMaximum dist. velocity in distillate channel 3.50 m/sDistillate corridor width 0.71 m

78 A.M. Helal et al. / Desalination 307 (2012) 76–86

with stages 16–18 making the heat rejection section. Connected tostage 18 is the de-aerator which receives the make-up flow.

The flash chamber shells are of welded rectangular construction15.8 m wide, with their length varying from 3.61 to 3.72 m in the heatrecovery section and from 3.9 m to 4.3 m in the heat rejection section.

The flash chamber shells are manufactured from carbon steel, exceptthepart above the demister in stages 5–18which are claddedwith Cu–Nito prevent corrosion. Parts not cladded with Cu–Ni are coated withepoxy coat. The water boxes connecting adjacent stages are alsomanufactured from carbon steel and cladded with Cu–Ni 90–10.

The evaporator is located on a reinforced concrete structure atan elevation of 4.3 m above the ground thereby providing enoughsuction head for various associated pumps.

Brine from the brine heater passes to the evaporator where a dis-tribution tube ensures even flow throughout the width of stage 1.Brine then flows through each stage with self regulation of the flowbeing achieved by means of adjustable gates on the orifices in eachinter-stage wall. The adjustment of gates can be carried out from in-side the evaporator. A sacrificial anode is secured to each brine gate.

A jump plate is located on floor of each stage (except stage nos. 1,9, 10 and 18). The function of the jump plate is to increase the waterdepth in front of the brine orifice thereby ensuring that the orifice isalways fully submerged.

A splash plate is fitted to the stage walls above the brine orifices toprevent brine splashes.

The re-circulating brine flows counter-currently through the con-denser tubes of stages 15 to 1 via the water boxes, while the seawaterpasses in a similar manner through stage 18 to 16 condenser tubes.

Suspended below the horizontal tube bundle in each stage is adistillate tray.

Vapor from the flashing brine passes through a demister, con-denses on the tube bundle with the droplets falling into the distillatetray and thence flows into a common channel (distillate corridor)located at one end, connecting the stages. From the distillate corridorat stage 18, a line leads to the distillate pump suction.

Each stage is provided with an internal baffle on the condenser tubebundle to provide a path for cooling the flow of non-condensable gassesbefore they are exhausted through the vent connections.

Various vent connections are provided from the stages to thevacuum system. There are various combinations in which thesevents will be used depending upon the operating conditions of theevaporator but under normal conditions the following will apply:

a. from stages 1, 2 and 3 separately to the first ejector condenserb. from stages 4–9 cascaded down via external stage connecting

pipes to the de-aeratorc. from stages 10–18 cascaded down via external stage connecting

pipes to the first stage ejector suction.

Vent and drain valves are fitted on each water box so that the aircan be expelled on start-up and the water boxes are drained whenshutting down.

The vessel will be drained from the last stage to discharge culvertvia the brine blowdown line. A rupture disk is provided on stage 9roof to protect the vessel from overpressurization.

3. MSF plant modifications to enhance productivity (themodified plant)

Fig. 2 is a schematic diagram for the modified plant, the subject ofthis work, which is characterized by two different distillate productlines, namely the recovery section distillate line derived from stage15 and the rejection section product line combining the individualcondensate streams from the three stages comprising the heat rejec-tion section. Unlike the interior design of the rejection stages in theconventional plant shown in Fig. 3, the distillate corridor has beenremoved in the modified plant design, Fig. 4. The distillate trayshave been replaced by longer ones extending below the total lengthof the condenser tube bundle and leading the stage condensate toflashing tanks where it is cooled before the collection in a main head-er and sent to the town water tanks by additional distillate pump. Inthe meantime, larger area demisters have replaced the old ones in therejection section, not shown in Fig. 4. Fig. 2, the modified plant dia-gram, shows the additional pumps over those normally found in theconventional plants on the condensate recycle line, the blow-downline, and product line in addition to the recovery section product line.

4. Literature review

In an early study, Tanios [3] investigated the capabilities of anexisting MSF plant to increase or decrease its capacity. He concludedthat the domain for such changes is limited. However, he applied hispredictions to an existing MSF plant at Shuwaikh, Kuwait, having adistillate capacity of 4546 m3/d (1 MIGD) and realized a 20% increasein production.

In an optimization study, Helal [4] investigated the feasibilityof uprating the three MSF distillers 4–6 at Umm Al Nar East‐AbuDhabi—UAE from 7.2 to 8.5 MIGD. The capacity was uprated whilemaximizing the gained output ratio (GOR) subject to all design andoperational constraints. Charts and tables have been produced forthe operation engineers to vary the plant capacity to any desiredvalue within the above mentioned range. Besides, the expectedperformance characteristics of the up-rated plants were tabulated. Itwas concluded that the main variables that contribute to the increaseof plant capacity are the top brine temperature and the recycle streamflow rate. The type of polymeric antiscalant used, proper maintenanceand the application of sponge ball cleaning system beside the properselection of materials of construction make it possible to uprate theexisting MSF plants within their over-design capabilities.

Awerbuch and Sommariva et. al. [1,2] discussed the capacityupgrading project of Sharjah Layyah desalination MSF unit 9. Plantupgrading is based on The Leading Edge Technology (LET) IntegratedHybrid Technologies which included the implementation of severalinnovative concepts in the plant design, as has been explained above.According to the authors, combination of those concepts contributedto the sensational results of addingmore than 45% distillate productionto the unit. The authors successfully introduced an innovative conceptwhich proved to significantly increase the plant production withminimal additional investment costs. Besides, modifications led toimproved purity in the last stages of MSF and increased flexibility ofthe plant to operate at increased capacity in winter without tempera-ture adjustment of inlet cooling water.

The great potential of nanofiltration membrane softening technol-ogy, developed by Hassan [5] and Al-Sofi et al. [6], was demonstratedby Awerbuch [1] through the first commercial nanofiltration system

Fig. 2. A schematic of the modified MSF desalination plant.

Demister

Brine pool

Orifice

Dis

tilla

te

corr

ido

r

Flashed off vapor

Demister

Brine pool

Orifice

Dis

tilla

te

corr

ido

r

Flashed off vapor

Tube Bundle

Distillate trough

Bundle

Distillate trough

Tube Tube

Fig. 3. Details of a conventional rejection stage.

79A.M. Helal et al. / Desalination 307 (2012) 76–86

Demister

Distillate trough

Tube

Demister

Brine pool

Orifice

Demister

Brine pool

Orifice

Flashed off vaporDistillate pipe to flash tank

Tube Bundle

Distillate trough

Fig. 4. Details of the modified rejection stage.

80 A.M. Helal et al. / Desalination 307 (2012) 76–86

integrated to increase capacity of existing MSF plant at Sharjah,Layyah desalination MSF unit 9, from nominal 5 MIGD to 7.2 MIGD(22,500 to 32,500 m3/d). This over 40% increase in capacity of MSFunit was a result of a two year demonstration and simulationprogram developed jointly with Sharjah Electricity and WaterAuthority, SEWA.

The additional capacity is achieved without building new intakestructure or new power plant in a very limited space which wouldnot allow construction of new desalination plant. The system involvesconstruction of NF plant to provide partial membrane softening offeed to MSF as well as modifications to existing MSF plant to becapable to achieve the increased capacity.

5. The process model

5.1. The general stage model (conventional MSF plant design)

Note: Equations are quoted here for a recovery stage, however, theequations are the same for the rejection stage with the substitution ofthe recycle flow rate inside condenser tubes in the recovery sectionwith the total intake feed rate in the rejection section.

For the rejection stages in the modified plant, the equation will bemodified accordingly.

5.1.1. Salt balance

Bj ¼ R � CR=CBjð1Þ

where:

Bj is the flashing brine mass flow rate off stage jR is the recycled brine mass flow rateCR is the recycle stream concentrationCBj

is the flashing brine concentration.

5.1.2. Overall mass balance

Dj ¼ R−Bj ð2Þ

5.1.3. Equilibrium correlation

TDj¼ TBj

−BPEj ð3Þ

where:

BPEj is the boiling point elevation evaluated at stage exit tempera-ture and concentration. Other parasitic losses are neglected,e.g. non-equilibration loss, temperature drops throughdemister and condenser.

5.1.4. Condenser enthalpy balance

TFjþ1¼ TFj

− Bj−1−Bj

� �� λj þ Dj−1 � CPDj−1

� TDj−1−TDj

� �h i= R � CPRjþ1

� �ð4Þ

where TFjis the recycle stream exit temperature from the stage

condenser.Note: the term [Dj−1∗CPDj−1∗(TDj−1−TDj

)], represents theamount of heat transferred at the condenser due to the re-flashingof the ingoing distillate to stage j. This amount of heat is used onlyto heat up the recycle stream inside the tubes along the length oftubes running across the distillate corridor and it does not help pro-ducing any portion of the stage product.

81A.M. Helal et al. / Desalination 307 (2012) 76–86

5.1.5. Heat transfer equation at the condenser

Uj ¼ R � CPRjþ1� TFj

−TFjþ1

� �= Aj � LMTD� �

ð5Þ

where the heat transfer coefficient Uj is taken as a constant value forone and the same stage and Aj represents the condenser surface areabased on the total length of the tubes, i.e., the length of the tubesexposed to the flashed off vapor plus the tube length running acrossthe distillate corridor. The log mean temperature difference, LMTD isgiven by:

LMTD ¼TFj

−TFjþ1

� �ln TDj

−TFjþ1

� �= TDj

−TFj

� �h i ð5� aÞ

5.1.6. Enthalpy balance on brine pool

TBj¼ TBj−1

−Vj � λj

Bj−1 � CPBj

" #ð6Þ

where

CPBj−1 is the specific heat capacity of flashing brine evaluated atthe brine temperature and concentration at the inlet fromstage number j and;

λj is the latent heat of evaporation of water evaluated at TBj .

For the case of modified rejection stages, Eq. (2) becomes:

Dj ¼ Bj−1−Bj ð2� aÞ

and Eq. (4) becomes:

TFjþ1¼ TFj

− Bj−1−Bj

� �� λj

h i.F � CPRjþ1

� �ð4� aÞ

where F represents the total feed flow rate, and Eq. (5) changes to;

Uj ¼ F � CPFjþ1� TFj

−TFjþ1

� �= Aj � LMTD� �

: ð5� aÞ

5.2.1. Brine heater enthalpy balance

Ws ¼ R � CPR � TBT−TF1ð Þλs

: ð7Þ

5.2.2. Brine heater heat transfer equation

UBH ¼ Ws � λs

ABH � LMTDBHand ð8Þ

LMTDBH ¼ TBT−TF1

ln Ts−TF1ð ÞTs−TBTð Þ

: ð8� aÞ

5.3. Blowdown splitter mass balance

R1 ¼ BN−BD ð9Þ

where BN is the brine rate off the last rejection stage, stage number N,and BD is the blowdown flow rate.

5.4. Makeup splitter mass balance

M ¼ F−WCW ð10Þ

where M represents the makeup flow rate and WCW is the rejectcoolant rate.

5.5.1. Mixer mass balance.

R ¼ M þ R1 ð11Þ

where R1 is the concentrated brine stream to be mixed with themakeup to form the recycle stream.

5.5.2. Mixer enthalpy balance.

M � CPM � TM−TBN

� �¼ R � CPR � TFNþ1

−TBN

� �ð12Þ

where TBNis the brine temperature off stage N which is taken as the

reference temperature for the enthalpy balance. N also representsthe total number of stages.

5.6. Physical property correlations

These functions can be referred to in Appendix A.

5.7. Model's assumptions

As it was decided to use the Excel Solver to solve the MSF plantmodels, and since a large number of iteration variables would berequired to be updated in every iteration, a condition that the Solvercannot handle, the following assumptions have been considered to bevalid in order to limit the number of iteration variables and enhanceconvergence to the final solution:

1. Concentration profiles in the heat recovery and heat rejectionsections of the plant were assumed to be linear, however, slopesof the two lines are different.

2. The flashing brine temperature profiles in the two plant sectionsare assumed to be linear. However, slopes (flash down per stage)of the two lines are not the same.

3. The overall heat transfer coefficient in each stage was calculated fromthe heat transfer equation of the specific stage disregarding the classi-cal correlation of the coefficient where its reciprocal represents thesummation of all resistances to heat transfer across the condensatefilm, the tube wall metal, the inside film and the scale deposit.

4. Temperature losses due to drops in vapor pressure across demistersand tube bundle have been neglected.

5. Assumption number 2 above makes Eq. (6) redundant. However,one of the criteria for convergence is to guarantee that; the maxi-mum absolute percent error between the value of the exit temper-ature of flashing brine,TBj , from the assumed linear profile and thatcalculated from Eq. (6) will be less than 0.7%.

6. Results and discussions

This work has two main objectives:

1- To investigate possibility of maximum productivity enhancementof the MSF plant, with the characteristics given in Table 1, keepingan eye on the operational constraints shown in Table 2 andsuggesting increase in pumping capacity whenever necessary.

Fig. 5. Plant capacity increase as a function of recycle rate at TBT=110 °C.

82 A.M. Helal et al. / Desalination 307 (2012) 76–86

2- To compare results for the conventional and modified plants inorder to find out the impact of the design modifications suggestedby Summariva [2] on the plant capacity enhancement.

As it has been mentioned before, the main contribution to capacityenhancement of the MSF plant comes through increasing the recyclerate and top brine temperature. This fact will be reconfirmed from theresults obtained in this work. In addition, we will emphasize theeffect of the design modification suggested by Summariva includingthe removal of the distillate corridor from the rejection sectionwhile diverting the accumulated distillate in the heat recovery sec-tion, after cooling, to the storage tank, thus preventing its successivere-flashing in the heat rejection stages.

In the first series of runs, and for each plant design, the recyclestream flow rate was increased step-wise from 0.725, the base case,to 0.85 of the maximum capacity of the recycle pump. These runswere performed maintaining the top brine temperature at 110 °C.To find out the effect of re-flashing inhibition of the distillate fromthe recovery section within the rejection section; the recycle rate,flash down per recovery stage and temperature rise across the brineheater, all were taken the same for both designs. This ensured samedistillate production rates from the heat recovery sections of theconventional and modified MSF plants.

Results in Table 3 show that for both plants, distillate capacityincreases steadily and linearly as shown in Fig. 5, with increasingrecycle rate, where for the conventional plant capacity, relative tothe base case conditions, capacity increase reaches 17.2% at a recyclerate 0.85* maximum recycle pump capacity. For the same recycle rate,percentage increase in production is 19.36% for the modified plant(Fig 6). The difference is only about 2% due to design modification.

In Table 4 we find that operating both plants at a recycle rate of0.85 of the maximum recycle pump capacity and increasing the topbrine temperature from 110 °C to 130 °C result in a total freshwatercapacity increase of 49.07% for the conventional design and 49.74%for the MSF plant with design modifications (Fig. 7). The reportedpercentage increase here is relative to the conventional plant capacityat the base case, i.e. when operated at a top brine temperature of110 °C and a recycle rate equal to 0.725×maximum recycle pump ca-pacity. The combined effect of brine recycle rate and top brine tem-perature on the product capacity of the conventional and modifiedplants is shown in Fig. 8.

With regard to the maximum pumping capacities listed in the lastcolumn of Table 5, it can be realized that except for the base caseoperating conditions, an auxiliary product pump will be required toassist the existing one in handling the added product capacity. Thesame argument is valid for the condensate recycle and blow-downpumps.

Table 3Effect of increasing recycle flow rate on plant capacity at TBT=110 °C.

Fraction of maxcapacity of recyclepump

0.725 0.75 0.775 0.8 0.85

Conventional design —

product capacity,m3/d

32,675 33,768 34,893 36,019 38,270

Modified design —

product capacity,m3/d

33,241 34,406 35,546 36,697 39,002

Conventional design%increase in plantcapacity⁎

0.00 3.35 6.79 10.23 17.12

Modified design, %increase in plantcapacity⁎

1.73 5.30 8.79 12.31 19.36

⁎ Relative to base case at 110 °C and 0.725×maximum capacity of recycle pump.

It is important to refer to contract that was signed by BESIX LeadingEdge Water Technologies, LET, and Mott Mac Donald, and SEWA, SharjaElectricity and Water Authority, to up-grade Unit 9 of Layyah powerplantwhich has a desaltedwater capacity of 5MIGD (22,500 m3/d) [1,2].

The addition of a nanofiltration unit for makeup softening up-stream the existing MSF plant at Layyah expectedly will enable re-moval of up to 99% of the SO4

−− ion, 50% of the Ca++ and Mg++

and HCO3− ions. Elimination of such scale forming ions permits safe

increase of the top brine temperature, TBT, up to 125 °C in the MSFplant, thus leading to an increase in capacity by 20–40%. The reducedhazard of scale formation helps extended operation with saving inchemicals cost. Meanwhile, modification of heat rejection stages ofthe MSF plant as described above results in improved performanceratio and reduction of the power to water ratio.

In the design phase of a new MSF plant, hybridization with an NFunit for makeup softening can lead to a number of advantages includ-ing; decrease in the design fouling factors and hence decrease in theheat transfer surfaces, and reduction in makeup flow due to salinityreduction. In addition, the reduction of non-condensable gassesreduces corrosion hazard.

Awerbuch [1] reported that annexing a nanofiltration unit for thesoftening of MSF plant makeup, and increasing recycle flow rate andtop brine temperature to about 125 °C increased theMSF plant capacityfrom 5 to 7.2 MIGD, i.e., an additional capacity of 44% could be realized.The additional cost of this hybridization is said to be less than thatrequired to build a separate RO unit, to work in parallel to the existingMSF distiller, with a capacity equal to the additional 2.2 MIGD.

Other attractive features regarding Layyah plant modifications canbe summarized in the following:

a- Minimum footprint where there is no room for a new desalinationplant

b- No changes to intake structurec- No increase in power facilities

Fig. 6. Percent increase in plant capacity relative to the base case, as a function of recyclerate (TBT=110 °C).

Fig. 8. Plant capacity increase relative to base case at 110 °C TBT and a recycle rate=0.725×maximum recycle pump capacity.

Table 4Effect of increasing top brine temperature on plant capacity at a recycle rate equal to0.85×maximum capacity of recycle pump.

Top brine temperature°C

110 °C 115 °C 120 °C 125 °C 130 °C

Conventional designproduct capacity —

m3/d

38,270 40,884 43,495 46,103 48,709

Modified design —

product capacity —

m3/d

39,002 41,312 43,770 46,383 48,926

Conventional design; %increase in plantcap.⁎

17.12 25.13 33.12 41.10 49.07

Modified design; %increase in plantcap.⁎

19.36 26.43 33.96 41.95 49.74

⁎ Relative to base case at 110 °C and 0.725×maximum capacity of recycle pump.

83A.M. Helal et al. / Desalination 307 (2012) 76–86

d- Cutting capital cost for additional capacity by 40%e- Reduction in operating cost.

On the other hand, the following limitations must be highlighted tohelp in making the right decision towards MSF capacity enhancement;

1) In the design of MSF plants, the recycle pump operation is normal-ly optimized so that; at steady state condition the pump is ratedfor the maximum energy efficiency. Increasing pump capacity toenhance the existing plant productivity will result in a lowerpump efficiency, hence a higher power consumption and higheroperating cost.Not only this, but also we got to know that long term operation ofthe recycle pump at a high discharge rate will increase the chancefor mechanical failures with the result of higher maintenance cost.

2) Higher pressure, higher temperature steam would be required atthe brine heater in order to heat up the recycle flow to a topbrine temperature of 130 °C. This will reduce the power plantcapacity in a cogeneration plant.

3) If the enthalpy associated with the product stream from the heatrecovery section is not utilized, then an additional heat exchangeror a flash tank and a condenser would be required to cool downthat stream before it is sent to the storage tank. This in fact is apart of the capital expenses of the modification project.

4) The addition of a new product line calls for additional productpump.

Table 6 compares the investment costs in $/(m3/d) for the twooptions of adding an NF unit to the existing MSF plant and buildinga separate RO unit to add 10,000 m3/d to the existing 22,500 m3/dcapacity of the MSF plant[1]. From the quoted figures it is seen that

Fig. 7. Plant capacity increase as a function of the top brine temperature, TBT, at aconstant recycle rate=0.85×maximum recycle pump capacity.

the specific investment cost per m3/d added is 755$/(m3/d) for theNF+MSF plant modification while it is 1110$/(m3/d) for the buildinga parallel RO plant. The saving in investment cost is (100∗(1110−755)/755) about 47%. In the meantime, the total cost from the modifiedMSF plant varies between 0.403 and 0.479 $/m3 versus 0.598 $/m3 forthe case of building an RO plant parallel to the existing MSF one. Thesefigures are given in Table 7.

For the specific case of Layyah MSF plant, Unit 9 modification,Table 8 reviews the investment costs while in Table 9 the total costper cubic meter of the additional capacity ranges between 0.403and 0.479 $/m3 as it has been mentioned above.

7. Conclusions

1- A simulation study has been conducted to investigate and comparethe capacity enhancement of a conventional brine recycle MSFplant and a modified one where the product from the recoverysection is diverted, after cooling, to the storage tank and where alter-ations have been introduced to the heat rejection section. Modifica-tion of the rejection section includes removal of the distillatecorridor, replacement of distillate trays and demisters. Both plantswere assumed to be coupled with a nanofiltration pretreatmentunit for makeup softening to enable increasing the top brine temper-ature to values up to 130 °C.The study was conducted using the operating conditions of Umm AlNar East 4–6 distillers. Abu Dhabi, The United Arab Emirates.

2- The impact of increasing the recycle flow rate up to 85% of themaximum capacity of the recycle pump, increasing top brinetemperature and modification of the rejection section, on plantcapacity increase were studied each apart.

3- It was found that the issue of rejection section design alterationalone contributes to minimal improvement in capacity enhance-ment where the modified design exceeded the conventional oneonly by about two percent increase in distillate capacity. Thisminor increase is predicted through plant modeling and in realityit could be negligible. Design modification for this specific portionof capacity increase is not recommended.

4- Increasing the recycle flow rate from 0.725 to 0.85 of the maxi-mum recycle pump capacity led to about 17.1 and 19.4% increasein plant capacity at a top brine temperature of 110 °C for theconventional and modified designs respectively.

5- At a recycle rate equal to 0.85 of the maximum recycle pumpcapacity, the increase in top brine temperature from 110 °C to130 °C boosted the total increase in plant capacity to reach avalue of 49.1 and 49.74% for the conventional and modified plantsrespectively.

6- In many cases of plant capacity enhancement, auxiliary pumpingmachinery need to be installed to assist the existing blowdown,condensate-recycle and product pumps. Higher temperature and

Table 5Effect of increasing the top brine temperature on MSF production capacity at maximum recycle rate (0.85×maximum recycle pump capacity).

TBT=110 Recycle rate=0.85×Rmax

Base case Conventional Modified Conventional Modified Max pump capacity Units

TBT=110 °C and 0.725×Rmax 0.85xRmax 0.85×Rmax 130 °C 130 °C

Dist rate 1,361,447 1,594,588 1,625,069 2,029,553 2,038,576 1,600,000 kg/hDist temp 37 38 36 38 37 °CRecycle rate 11,172,975 13,099,350 13,099,350 13,099,350 13,099,350 15,411,000 kg/hRecycle concentration 68,000 68,000 59,703 68,000 60,739 ppmRecycle temp 39.9 39.9 39.5 40.3 40.6 °CMakeup rate 5,091,806 5,981,826 9,904,303 7,613,521 11,453,265 kg/hMakeup concentration 50,000 50,000 50,000 50,000 50,000 ppmMakeup temp 43.2 43.3 42.8 43.3 43.3 °CFeed rate 11,612,274 13,711,998 13,093,021 14,113,025 12,577,449 16,775,000 kg/hFeed concentration 50,000 50,000 50,000 50,000 50,000 ppmFeed temp 32 32 32 32 32 °CBD rate 3,743,975 4,398,401 8,294,699 5,598,178 9,428,202 5,555,000 kg/hBD concentration 68,000 68,000 59,703 68,000 60,739 ppmBD temp 39.9 39.9 39.5 40.3 40.6 °CReject coolant rate 6,520,469 7,730,172 3,188,718 6,499,503 1,124,184 kg/hGOR 6.36 6.35 6.47 7.80 7.84 kg dist/kgstm.Top brine temperature 110 110 110 130 130 °CSteam temperature 120.0 120.0 120.0 140.0 140.0 °CSteam rate to the brine heater 214,225 251,161 251,161 260,115 260,115 220,000 kg/hVapor velocity at last rej. stage 6.75 7.87 9.20 7.80 8.39 m/sVapor velocity at 1st rec. stage 0.59 0.69 0.69 0.53 0.53 m/sRecycle velocity in tubes 1.55 1.81 1.83 1.81 1.82 m/sSeawater velocity in tubes-rej. sec. 1.71 2.02 1.93 2.08 1.85 m/sDist. rate in the base case 1,361,447 1,361,447 1,361,447 1,361,447 1,361,447 kg/h% increase from the base case 0.00 17.12 19.36 49.07 49.74 %

Underlined values means that the maximum capacity of the corresponding pump will be exceeded and additional pumping capacity will be required.

Table 6Investment cost comparison RO vs. NF+MSF to add 10,000 m3/d to the existing22,500 m3/d capacity of MSF plant [Reference [1]].

$ INV/(m3/d) RO NF+MSF

MSF modifications N/A 111NF trains N/A 510.6Membranes 732.6 33.3Civil work 199.8 55.5Consulting/engineering 222 22.2Misc. and commercial 22.2 22.2Total direct investment 954.6 754.8Intake 155.4 N/ATotal investment 1110 754.8

84 A.M. Helal et al. / Desalination 307 (2012) 76–86

pressure steam is required which deprives the steam power plant(in a dual purpose plant) some of its power production. Also, anadditional heat exchanger or flashing tanks to cool the recoverysection product stream before being sent to the town water tank(in the case of modified design) is required.

7- Cost analysis of other workers showed that the specific cost percubic meter of the additional capacity obtained as a result of theaddition of an NF softening unit upstream an existing MSF plant,TBT and recycle rate increase, would be less than the unit costresulting from building an RO plant with the same additional ca-pacity and working parallel to the MSF plant.

Table 7Unit product cost analysis RO vs. NF+MSF [1].

RO $/m3 NF+MSF $/m3

Annual investment return⁎ 0.330 0.177–0.245Operating costs 0.268 0.226–0.234Total 0.598 0.403–0.479

⁎ Based on a 20 year life and 6% interest rate.

The addition of an NF unit for partial purification of the makeupstream to enable a scale-free operation and to go up with the topbrine temperature upsets design simplicity of the traditional MSFplant and complicates its process control more.

Nano-filtration of the Gulf water with its high content oforganic carbon may call for a strict pretreatment scheme to avoidbio-fouling of membranes which adds to the capital and operatingcosts (this limitation applies to RO membranes as well). In addition,a larger portion than 25% of the makeup stream may require to bepurified to ensure a scale-free operation which means a larger NFunit and additional capital cost which may defeat the economicfeasibility of plant modification.

Until now, economic feasibility of adding an NF unit to an existingMSF plant to enhance the capacity of the latter is not well proven andenough cost data is not available to make the right decision.

In general, comparison between the advantages and disadvan-tages of increasing the existing capacity of an MSF plant beside thelack of information about operational problems, makes it hard totake a decision whether to go for MSF plant capacity increase orbuild a new RO unit of the same added capacity. However the absenceof enough ground area in the vicinity of the existing MSF plant andthe need for additional water capacity may favor the option of MSFplant modification.

Table 8SEWA investment cost for 10,000 m3/d additional capacity [Reference [1]].

Costs $ [multiply by 1000] $ invested/(m3/d)

MSF modifications 1108 111NF trains 4616 510.6Membranes 343 33.3Civil work 328 55.5Consulting/engineering 222 22.2Misc. and commercial 643 22.2Total 7402 754.8

Table 9SEWA — additional unit cost analysis [Reference [1]].

$/m3

Annual investment return⁎ 0.177–0.245Operating cost 0.226–0.234Total 0.403–0.479

⁎ Based on a 20 year life and 6% interest rate.

85A.M. Helal et al. / Desalination 307 (2012) 76–86

8. Symbol list

A Heat transfer area, m2

B flashing brine mass flow rate, kg/hBD Blowdown mass flow rate, kg/hBPE boiling point elevation, KCB Flashing brine concentration, kg salt/kg solutionCPB, CPD, CPF, CPM, CPR Heat capacities of flashing brine, distillate,

seawater feed, makeup and recycle streams respectively,kJ/kg.K

CR recycle stream concentration, kg salt/kg solutionD Distillate mass flow rate off stage jLMTD Log mean temperature difference, KM Makeup mass flow rate, kg/hR Recycle brine mass flow rate, kg/hR1 Mass rate of concentrated brine stream mixed with the

makeup to form the recycle stream, kg/hTD Distillate temperature, KTB Flashing brine temperature, KTF Recycle stream exit temperature from the stage condenser, KTBT Top brine temperature, KU Overall heat transfer coefficient, kW/m2.KV Mass flow rate of flashed off vapor at stage j, kg/hWs Mass flow rate of saturated steam to the brine heater, kg/hWCW Reject coolant mass rate, kg/h

SubscriptsBH Brine heaterj Stage numberN Total number of stages.S Steam to brine heater

Greek lettersλ latent heat of evaporation, kJ/kg

Appendix A. Physical and thermodynamic property correlations

N.B. All correlations given below are applicable for ranges safelycovering all temperatures and salinities reported in this paper.

A.1. Pressure of saturated steam as function of temperature, [7]

P ¼ 101:32 � 218:167 ��10

�̂�−TT=TK � ð3:2437814þ 0:00586826

� TT þ 0:000000011702379 � TT3�= 1þ 0:0021878462 � TT

����

P saturation pressure, kPaTT Tc–TKTc Critical temperature of water=647.27 CTK Saturation temperature, K

A.2. Specific heat capacity of purewater as a function of boiling temperature,[7]

CpD ¼ 4:1868 ��1:0011833−0:000061666652 � TF

þ 0:00000013999989 � T2F þ 0:0000000013333336 � T3F�

��0:011311−0:00001146 � TF

�

CpD Specific heat capacity of water kJ/kg.°CTF Boiling point of water °F.

A.3. Heat capacity of brine as a function of temperature and concentration[7]

CB ¼�4:1868 �

�1:0011833−0:000061666652 � TF

þ0:00000013999989 � T2F þ 0:0000000013333336 � T3F��

��1−

�CB � 0:000001

���0:011311−0:00001146 � TF

��

Where

CpB Specific heat capacity of brine solution, kJ/kg.KTF Brine temperature, °F.CB Salt concentration, ppm

A.4. Density of brine solutions as a function of temperature and saltconcentration, [7]

ρ ¼ 16:0034 ��62:707172þ 49:364088 �

�CB � 0:000001

�−0:0043955304 � TF−0:032554667 �

�CB � 0:000001

��TF−0:000046076921 � T2F þ 0:000063240299

��CB � 0:000001

�� T2F

�

ρ Density, kg/m3

TF Brine temperature,°F.CB Salt concentration, ppm.

A.5. Latent heat of vaporization of water as function of the boilingtemperature, [8]

λS ¼ 2501:9−2:407 � Tþ 0:0011922 � T2−0:000015863 � T3

T Saturation temperature, CλS Latent heat of vaporization, kJ/kg

A.6. Specific vol. of vapor, [8]

v ¼ 0:003172222 ��647:286=TK−1

��EXP

�83:63213098−0:668265339 � TK þ 0:002495964 � T2K

−0:00000504185 � T3K þ 0:00000000534205

�T4K−0:0000000000023279 � T5K�

v Specific volume of saturated vapor, m3/kgTK Saturation temperature,

86 A.M. Helal et al. / Desalination 307 (2012) 76–86

A.7. Boiling point elevation as a function of temperature and saltconcentration, [7]

BPE ¼��

CB � 0:000001�� T2

K=13832���1þ 0:001373

� TK−0:00272 ���

CB � 0:000001�̂

2�� TK þ 17:86 �

�CB

� 0:000001�−0:0152 �

�CB � 0:000001

�� TK

��TK−225:9

�=�TK−236

��−�2583 �

�CB � 0:000001

���1−

�CB

� 0:000001��

=TK

�

BPE Boiling point elevation, CTK Brine temperature, KCB Brine concentration, ppm

References

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[2] C. Sommariva, L. Awerbuch, Improving efficiencies in thermal desalination systems:the Layyah plant experience, in: IDA International Water Forum-Innovation andIntegration — Impact on Desalination and Water Reuse Costs, Grand HyattConvention Centre— Dubai, UAE March 5–6, 2006, 2006.

[3] B.Z. Tanios, Marginal operation field of existing MSF distillation plants, Desalination51 (1984) 201–212.

[4] A.M. Helal, Uprating of Umm Al-Nar 4–6 MSF desalination plants, Desalination 159(2003) 43–60.

[5] A.M. Hassn, Development of a Novel NF — seawater desalination process andreview of application from pilot plant to commercial production plant stages, in:The Third International Water Conference in the Arab Countries, Beirut— Lebanon,27–30 September, 2004 (ARWATEX — 2004), 2004.

[6] M.A.K. Al-Sofi, A.M. Hassan, G.M. Mustafa, AGh. Davi, N.M. Kither, Nanofiltration asmeans of achieving higher TBT > 120oC in MSF, in: Conference for Membranes inDrinking and IndustrialWater Production, Amsterdam, 21–24 September 1998, 1998.

[7] A. Helal, M.S. Medani, M.A. Soliman, J.R. Flower, A tri-diagonal matrix model formultistage flash desalination plants, Comput. Chem. Eng. 10 (4) (1986) 327–342.

[8] H.M. Ettouney, H. El-Dessouky, F. Al-Juwayhel, Performance of the Once ThroughMultstage Flash Desalination Process, Proceedings of the Institution of MechanicalEngineers. Part A. J. Power Energy 216 (3) (2002) 229–241.