Process modelling in desalination plant operations

Desalination 222 (2008) 431–440

Presented at the conference on Desalination and the Environment. Sponsored by the European Desalination Societyand Center for Research and Technology Hellas (CERTH), Sani Resort, Halkidiki, Greece, April 22–25, 2007.

Process modelling in desalination plant operations

Walid ElMoudir*, Mohamed ElBousiffi, Salah Al-Hengari Libyan Petroleum Institute, P.O. Box 6431, Tripoli, Libya

Tel. +21892 5227107; Fax +21821 4830031; email: [email protected]

Received 5 January 2007; accepted 10 January 2007

Abstract

Operation of desalination process is sometimes a difficult task, especially when the plant is getting older.Process engineers and operators are required to follow-up any changes or variations in process parameters eitherin control room or on-site measurements.

Undetectable deviation and/or malfunctioning in any of measurement devices could occur specially duringlong-time successful operation without major shutdown for the entire plant or when the process is part of verycomplex integrated processes in an industrial site. This deviation, inaccurate-reading or malfunctioning values canbe due to un-calibrated devices (i.e. flow-meters) or fouling of measurement devices (i.e. thermo-couples).

Practically, process modelling is a useful tool that can assure and confirm that the process measurementsparameters are accurate. Incorrect balancing of mass and/or energy can be detected easily as well as the productivityof the plant. In this paper, multi stage flash (MSF) desalination plant was illustrated and evaluated. Malfunction ofmeasurement device showed the plant was producing distillate more than what was really measured or received bydown stream processes. Excel spreadsheets was developed and utilized for process modelling and confirming theresults.

Keywords: MSF desalination process; Process modelling; Plant data; Maloperation

1. Introduction

Desalination is used whenever there is ashortage of water from natural resources. Inindustrial sites, high quality water is essentiallyneeded for production of steam or for possessingpurposes. Many desalination technologies findtheir way for industrial application and this

includes multi stage flashing process (MSF). Thisprocess has gained good reputation for years dueto its high productivity capability, operability andflexibility (110–70% design capacity [1,4]).MSF consumes energy from low grade sources(i.e. LP steam generated by waste steam genera-tor). In many regions in the world, good operat-ing experiences are gained from running thistechnology, for years. *Corresponding author.

0011-9164/06/$– See front matter © 2006 Published by Elsevier B.V.doi:10.1016/j.desal.2007.01.170

432 W. ElMoudir et al. / Desalination 222 (2008) 431–440

The desalination operation is a critical importanceto operation of other chemical and petrochemi-cal plants within a chemical complex or refin-ery. Failures are experienced and plant forceshutdown usually occurs leading to considerableaffects or effects on the plant economics. Thedesalination plant economics can be affected bythe plant productivity and availability (Fig. 1).The productivity and availability factors areused to describe the production capacity and howlong the plant is running during a specific periodof time (i.e. a year). Combined production andavailability factors formulates the plant factor(plant factor = production factor × availabilityfactor) which is important indicator of plant per-formance.

This paper illustrates some operating experi-ence of MSF plant. The plant had maloperationcase occurred during normal operation period.Process modelling was used to detect, determineand confirm the maloperation case.

1.1. MSF-BR process description

All MSF processes are widely utilized; theyare the driving force for desalination. The per-centage of MSF installed capacity over the totaldesalination installed capacity world wide isover 50% [3]. The objective of the MSF system

is to overcome the main drawback of the singlestage flash unit that is to improve the system per-formance ratio; by another word, improve theproductivity and minimize energy consumption.This is achieved by dividing the flash range overa larger number of stages.

MSF commercial systems are divided into twoprocess configurations; multi stage flash oncethrough (MSF-OT) and multi stage flash withbrine circulation (MSF-BR). From their names,in the first process feed seawater passes throughout the process once through at a time while in thelater process a small amount of seawater feed isused and mixed (Fig. 2) with major recyclingflow of concentrated flashing brine exist fromthe last stage. In this way, control of salt concen-tration is achieved with less fouling problems.MSF-BR has considerable attention due to manyadvantages over the MSF-OT such as less waterfeed (make-up), less chemical consumptions andtreatments, and higher performance ratio [3].

The energy is supplied to the brine heater andthen it is removed in the heat rejection section.All of the heat recovery stages are identical inconstruction as well as heat rejection section.The feed seawater is drawn through the intakepipes, pre-treated, pre-chlorinated and suspendedsolids removal by filtration. Cold seawater entersthe condenser tubes in the last stage (stage 16)

Fig. 1. Influence of plant factor on plant economics [2,12].

W. ElMoudir et al. / Desalination 222 (2008) 431–440 433

of the process at an initial temperature of seawa-ter. Then warm seawater feed coming out fromheat rejection section is divided into twostreams. About 40% of warm seawater leavingthe rejection section is discharged to the effluentline and about 60% of this warm water is divertedas a make — up feed to the evaporator in orderto maintain a constant salt concentration incirculating brine in the evaporator.

The brine circulation is heated by condensedwater vapor in the condenser/vapor side until itexits at a hot temperature. Then it is heated to topbrine temperature in the brine heater by LP steamand enters the first stage flashing chamber. Aftersuccessive flashing, the brine finally leaves theunit at concentration below 70,000 ppm. The dif-ference between top brine temperature (TBT)and brine blowdown temperature is called theflash range which is the driving force of multistage flashing processes.

Dissolved CO2 and air must be removed (bydeaerator and venting ejectors) from seawatermake-up stream to keep CO2 concentration at avery low level, in order to prevent corrosion, andto minimize the amount of non-condensable gaseswhich would impair the heat transfer rate. Anti-scale chemicals are injected to the seawater makeup stream in order to eliminate scale formation.

2. Case study: MSF-BR plant

MSF-BR plant was considered in this paper.The plant was fabricated from good corrosion-resistant material; it has been in operation formore than 20 years. The plant was designed todeal with Mediterranean Seawater conditions. Itcould run in two operation modes. The first modeis operating at high top brine temperature (TBT =115°C) with production capacity of 8400 t/day(350 t/h). The second mode is operating at lowerTBT of 90°C with production of 6000 t/day(250 t/h). As seawater temperature seasonalchange (Table 1), the plant could change its

Fig. 2. MSF-BR desalination process.

Table 1Main plant specification data

aPerformance ratio = distillate production/steamconsumption.

Items Units Operation condition

Case 1 Case 2 Case 3

Seawater feed temp. °C 16 24 27 Maximum brine temp. °C 90 90 90 Distillate production t/h 277 250 240 Steam consumption t/h 44.25 40.00 38.52Performance ratioa – 6.26 6.25 6.23 Distillate purity (TDS) ppm <3.0 <3.0 <3.0


operating conditions. Production can beincreased to 277 t/h during cold season as thedriving force (flash range) increases althoughthe steam consumption would be increased too.The opposite could occur during summer hotseason (seawater temperature ≥ 27°C).

Although, the plant was successfully commis-sioned to run at both operation modes, the waterdemand is satisfied by running the plant at topbrine temperature of 90°C. In some events ofwater shortage and/or increase of water demands,operation at higher temperature has occurred butnot at 115°C.

However, from practical experience theproblems were experienced when the plant oper-ates at designed high top brine temperature(TBT = 115°C). Fouling formation rate was sig-nificantly increased inside the plant heatexchangers’ tubes and ultimately led to forcedshutdown for cleaning. More details are pre-sented in the following section.

Therefore, the plant is kept to run at productionrate of 250 t/h (TBT = 90°C). In the recentyears, the normal production was below nominaldesign value (250 t/h) due to many reasons asthe plant getting older. Some important com-ments and observations on the plant situation arenoticed: (1) The plant suffers from corrosion attacks

internally and externally, (2) Some measurements devices were out of

service and some other are not calibrated, (3) Some leaks and vacuum losses were experi-

enced due to corrosion problems forcing theplant operators to operate the venting sys-tem at full load to control the stages’ pres-sures inside the plant. Although, it is notrecommended as it is affecting the ventingsystem. It leads to losing some of the pro-duction potential in steam which will bedrawn from production vapor space in eachstage as well as the motive HP steam used inventing system.

3. Fouling problems

In MSF process, anti fouling chemicals areused to inhibit crystallisation fouling formation.Three types of foulants are expected to form inMSF process. They are usually deposited becauseof the reverse solubility of salts present in seawaterlike CaCO3, Mg(OH)2 and CaSO4 when the tem-perature is increased. The anti-fouling chemicalsare used to mitigate or minimize formation ofalkaline salts (i.e. CaCO3, Mg(OH)2) and thesechemicals are used and showed very effectiveresults in the plant under investigation. However,antiscale chemical used in the plant and accord-ing to design manufacturing company was basedon polyphosphate anti-foulants at first years ofoperations. This chemical is known to decom-pose easily at high temperatures (>100°C) leadingto loss its efficiency. In recent years, a new effec-tive and stable anti-foulant was used and prove tobe effective on CaCO3 and Mg(OH)2 scales.

However, non-alkaline salts (i.e. CaSO4 salts)are only controlled through both brine salinity andtemperature control. Therefore, the MSF plantsare designed to operate away from CaSO4 saltsformation region. The maximum brine blowdownconcentration is limited by 70,000 ppm (designcriteria of MSF [3]) and the maximum tempera-ture recently reported to be 122°C [9] when effec-tive anti-foulants are used, i.e. Belgard EVN.

Brine salinity of 70,000 ppm refers to the max-imum allowable concentration where beyond it;CaSO4 salts solubility might be reduced leadingto enhance fouling problems. Fig. 3 illustratesthe solubility curves of three salts of CaSO4

which the most likely to precipitate. The plant isaway from problems of exceeding solubility ofnearest salt (anhydrite) when the plant operatesat 90°C (Case 1, in Table 1).

It was mentioned in the literatures that thenucleus of anhydrite CaSO4 need long period oftime to start to form crystals and precipitate. Asthe fabricated company of the plant claimed, theCaSO4 salts will not reach saturation point and


time fo brine to pass through brine heater andheat exchangers is very small in any operationscenario. Their claims are supported with whathave been found in literatures [7–9]. Therefore,there is no fear from precipitation problems at all.

In practice and according to the plant operators,this is not fully accurate statement. The plantsufered from severe fouling when it was opera-ting at TBT = 115°C. Photo 1 shows the cross sec-tion view of its brine heater when the plantoperated for a period of time at TBT = 115°C. Thescale was concentrated in hot outlet location of

two-pass flow brine heater. Tube blocking ofheat exchanger outlet was more than 50% whileinlet tubes were clean with no blockage. Sincethen, the plant operates at or close to TBT =90°C, with no fouling problem reported.

Fig. 4 explains the reasons behind the foulingthat was experienced at TBT = 115°C. The recyclebrine was entering to the brine heater at CaSO4

anhydrite super-saturation region; therefore, thefouling was occurred.

4. Plant operational history

The plant has a successful operation historyduring last 20 years of operation. Availabilityof experienced operators personal is importantreason of keeping the performance of this plantat acceptable levels. However, it was reportedthat there were huge mismatching in the plantdata recorded during recent operating months.

The plant production was not matching theamount of water received either by steam plantand/or in water storage tank farm. This raised somedoubtful about plant data that being recorded byboth control room and measurement devices inthe site. Productions and consumptions were notin homogeneity as material and energy balances

Fig. 3. CaSO4 solubility curves (Tsea = 24°C & TBT =90°C) [12].

Photo 1. Fouling in brine heater after operation for aperiod at TBT = 115°C.

Fig. 4. CaSO4 solubility curves (Tsea = 24°C & TBT =115°C) [12].


show and maloperation. Fig. 5 illustrates theproduction and steam consumption of the recentnine months of operation. These values wereobtained from the technical department as thebest estimation of plant production. However,these data were different from recorded at thesite. The technical department was noticed thatthese data was not matching the data which wasreported of the operation department and therewas lack of homogeneity between recorded andestimated production rates and consumptions.

A troubleshooting plan was put in action.Selected and scattered daily Log-sheets werecollected and evaluated. Production and con-sumption rates of different plant streams wereevaluated. Table 2 shows Plant daily log-sheetdata along with reference data.

Daily plant data are collected every two hoursfrom measurements instruments in the site. Infirst column of Table 2 reference data of plantwhen it produces 250 t/h while the second columnshows the reading from measurements instru-ments. It was recorded that the production

capacity was 200 t/h. It is worthy mentioningthat the flow-rates of seawater feed, make-up, con-densate, and recycle brine were all correct withaverage standard deviation of less than 5%.

4.1. Review of plant operational conditions

Production of MSF process can be estimatedby two applicable methods, the simple methodand detailed stage-by-stage method. Whether thecalculations are for designing case or operationcase, both methods are valid and reliable.

4.2. Production estimation method

It is a preliminary method to estimate theproductivity of MSF-BR unit. Further details onthis method are described elsewhere [5,11]. Thismethod goes back to the principle of MSF pro-cess, it allows carrying out general diagnoses ofmaterial and energy balances.

The concept of this method is based on thetemperature flashing range (TBT – Brineblowdown temp.) which is a driving force of

Fig. 5. Estimation of distillate production and steam consumption (9 months).


MSF production. Accordingly, the MSF thermalefficiency of the unit is dependent on the tem-perature range. This method assumes ideal sit-uation and ignores all types of thermodynamicsand pressure drop limitations such as boilingpoint elevation (BPE), non-equilibrium allow-ance loss (NEA), temperature and pressure lossesin demisters and tube bundles, and heat loss fromthe process to environment. The method in itsexisting form seems applicable to MSF-OT pro-cess; however, it can be applied for MSF-BR astheir design quite close to each other.

Fig. 6 shows a schematic diagram of MSF-BR process with simple method input parame-ters. It makes possible to estimate productivityof plant from few plant data and provide estima-tion of heat load supplied to the plant. It canhelp to identify the wrong data shown in themeasurement instruments regarding productionif there is any.

Product water = Md = {Mr × Cp × Flash range}/λ (1)

Heat input or brine heater load = Qin = (((T3 – T2) × λ)/Flash range) × Md (2)

where T1 = inlet seawater temperature (°C) T2 = hot seawater temperature inlet to brine

heater (°C) T3 = top brine temperature (°C) T4 = brine blowdown temperature (°C) Flash range = T3 – T4 (°C) Mr = flashing brine flow rate (t/h) Cp = average specific heat of flashing brine

across evaporator (kJ/kg °C) λ = average latent heat of vaporization across

evaporator (kJ/kg)

From data given in Table 2 for the referencecase, (the simple method)

Table 2Daily log-sheet of plant on certain day

aReference case is operation case when seawater temperature is 24°C.

Stream Units Referencea Control room

Brine heater Brine temp. in/out °C 82.7/90 80/88 Heating Steam temp. °C 95 104 Condensate t/h 39.792 34

°C 95 103 μs/cm <1 1.7

Heat rejection section Distillate t/h 250 200 °C 37.9 Out of order μs/cm 3 ppm 1

Brine blowdown t/h 364.37 300 °C 39.2 42

Seawater make-up t/h 614.37 500 °C 38.3 40

Seawater inlet t/h 1505 1200 °C 24 25 μs/cm 42,180 ppm 40,000

Brine recycle t/h 3098 2400 °C 39 39 μs/cm 65,380 ppm Out of order


Md = 259 t/h (total production of the unit) Qin = 24,248 kW (total heat input to the unit)

Total production (Md) is 259 t/h which isclose to the design production (250 t/h). Thedifference is 9 t/h, which is 4.5% higher thandesign value, sum of all types of losses occursin the plant. For heat input, the heat load (Qin)is 24,248 kW while the design load is24,507 kW. The difference (258.5 kW), which is1% higher than design value, is very small and atacceptable level.

At this stage and in the same manner of pre-vious calculations, this simple method can beused to check the plant productivity and heatconsumption of real operating data mentioned inTable 2. Md = 185 t/h (total production of the unit) Qin = 20,801 kW (total heat input to the unit)

The production is less than 185 t/h. Therefore,it easily recognizes recorded production valueinside plant was inaccurate in the Table 2. Thetypical (ideal) case of production will not exceed185 t/h which ignores all type of loss.

4.3. Process modelling

Desalination plants are designed based onmaterial and energy balances. It is possible todevelop a mathematical model to describe such

balances. El-Dessouky and Ettouney developedand demonstrated simple and reliable models formany thermal desalination processes in theirrecent text-book [3]. These mathematical modelscan be easily programmed in Fortran, MatLabor even in Excel Spreadsheet combined withVB programming language.

Process modelling is a tool used to check plantoperation through determination material andenergy balances. It provides significant look atparameters and variables which might not possibleto be available for plant operators or can not bemeasured with measurements devices. Utilizingthe process modelling for MSF-BR was imple-mented and confirmed some unreliability ofplant data.

4.3.1. Utilizing of MSF excel spreadsheet

The Libyan Petroleum Institute (LPI) hascarried out a study to evaluate the performance ofsome desalination plants in the oil sector. Duringthe course of this study, a number of excel spread-sheets were developed among them are twofor modelling of MSF-BR (LPI_MSF-BR) andMSF-OT (LPI_MSF-OT). The spreadsheetsinclude material and energy balances calculationsas well as heat transfer characteristic of theprocess. The mathematical models are based onthe El-Dessouky and Ettouney Models for MSFprocesses [3,6] and some correlations from

Fig. 6. MSF-BR with input parameters of simple method.


another ref. [11]. The model considers lossessuch as thermodynamic losses or pressure droplosses in condenser and mist eliminators.

The philosophy of these Excel spreadsheetsis that through a user supply data for (i.e.streams flow rate and temperature) of a unit, thecalculation will processed from stage-to-stage untilthe outlet streams. The calculations start fromHot section (i.e. Brine Heater) in the unit up toits outlet streams (i.e. distillate and brine blow-down). The results are tabulated and plotted.

The plant data (Table 2) will be examined byMSF modelling built in Excel spreadsheet. Theestimated production of the plant was 182 t/h(Fig. 7). The spreadsheet estimation is moreaccurate than the preliminary method shownpreviously. The recorded production of the plantwas inaccurate (200 t/h).

The corrective action was taken to fix the mal-functioning of measurement device and the realcapacity was matching the simulation results.Maintenance team from the operating companywas investigating measurement instruments.Calibration, fixing and re-setting for all instru-ments and pumps were carried out. Examiningvalve performance and replacements of some

instruments with proper type were done. Sincethen and on regular basis, the operating com-pany operators carry out check on the plant massand energy balances and comparing the resultsagainst recorded values.

5. Recommendations

(1) It is vital for the desalination plants to havea routine checking scheduled maintenanceprogram in order to guarantee an efficientand reliable units,

(2) Quick measures should always be takenin case of unforeseen developments duringoperation in order to avoid costly downtimeand loss of production,

(3) Process modelling for desalination processesis an important tool that can help to identifythe main process parameters and variables,

(4) Process software modelling is recommendedto be used on daily basis in order to allocatingwrong data easily and quickly,

(5) Process modelling (steady state and dynamic)is recommended to be used to check differentoperational scenarios and strategies beforeapply on the plant,

Fig. 7. Design conditions at seawater temperature is 25°C (LPI_MSF-BR).


(6) LPI MSF spreadsheets can be used to checkthe operation of MSF plants. It can be usedas a training tool for new engineers andtechnicians in these plants,

(7) Attention should be paid to the fouling fac-tor since this playing a critical rule in thecalculation of heat transfer; therefore, thereis a need to improve and review old foulingfactor values [10] used in designing of olddesalination plants,

(8) It is recommended to modify the processconfiguration of MSF-BR by fixing seawa-ter temperature fed to the plant in order tostabilize operation and production. This canbe achieved by mixing the cooling seawaterwith seawater intake.

Acknowledgement

Authors would like to thank the LibyanPetroleum Institute for supporting this work andpermission to publish the data and results.

References

[1] S. Al-Hengari, M. El-Bousiffi and W. El-Moudir,Libyan Petroleum Institute experience in evaluationof desalination plants in the Libyan oil sector,Desalination, 206 (2007) 633–652.

[2] W. El-Moudir, M. El-Bousiffi and S. Al-Hengari,Economic Analysis for Thermal Desalination

Processes, Technology of Oil and Gas Exhibi-tion and Forum 2004, Tripoli, Libya, 2004.

[3] H. El-Dessouky and H. Ettouney, Fundamental ofSalt Water Desalination (1st edn.), Elsevier Sci-ence B.V., Amsterdam, 2002.

[4] M. Darwish, H. El-Dessouky and H. Fath, ShortTraining Course on Desalination Technologies,Petroleum Research Centre, Tripoli, Libya, 19–31January 2002.

[5] www.rpi.edu/dept/chem-eng/boitech-environ/environmental/desal/nuchight.html

[6] H. El-Dessouky, Steady State Analysis of MSFDesalination Process, Desalination, 103 (1995)271–287.

[7] M.A. Darwish, M. AlSaied, A. AlSaied andS. Ali, Engineering Systems of Seawater Desali-nation (Arabic Edition), Scientific PublishingCenter at the University of King Abdul Azziz,Jeddah, Kingdom of Saudi Arabia, 1995.

[8] H. Fath, Short Training Course on Operation andMaintenance of Desalination Processes, Petro-leum Research Centre, Tripoli, Libya, 12–16 July2003.

[9] Mohammad Abdul Kareem Al-Sofi, Fouling Phe-nomena in Multi Stage Flash (MSF) Distillers,Desalination, 126 (1999) 61–76.

[10] Standards of Tubular Exchanger ManufacturesAssociation, TEMA, Inc. New York, 1979.

[11] A.H. Khan, Desalination Processes and Multi-stage Flash Distillation Practice, Vol. 1, Elsevier,Amsterdam, 1986.

[12] French Oil and Gas Industry Association, SeawaterCircuits: Treatment and Materials (Translation fromFrench Lang.), Editions Technip, Paris, 1998.

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Process modelling in desalination plant operations