don't let the pinch

It has now been more than 20 yearssince the pioneering work of Umedaand his coworkers at Chiyoda laid thefoundations of modern process inte-gration technology (1). Over the four-year period of 19781982, this team in-troduced the concept of overall processsystem analysis, reintroduced the conceptof composite curves and showed howthese could be used to determine the utili-ty needs of a process, developed a syn-thesis strategy based upon overall net-work costs, identified the heat recoverypinch for the first time, demonstratedhow design of the core process and theheat-recovery system could be linkedthrough manipulation of streams aroundthe pinch point, and introduced the heatdemand and supply (HDS) diagram (nowmore commonly referred to as the grandcomposite curve) and the raffinate dia-gram (now used in overall site analysis)(14). An overview of the technology isnow available on the Internet (atwww.Pinchtechnology.com).

The methodology was subsequentlyextended and popularized, most notablythrough the development of the pinchdesign method (which moved networksynthesis away from the computer andinto the hands of practical engineers) andthrough the introduction of the powerfulconcept of target before design. Withthese later developments, the processpinch became the pivotal point in heat-recovery network design, and the pinchhas now become the focus of the method-ology. (The method has even becomeknown as Pinch Technology.)

The Chiyoda team did not place asmuch emphasis on the pinch as is donetoday. To them, the pinch was importantin directing process modification ratherthan controlling design. They had validreasons for this philosophy. They ob-served that, although it is possible to con-struct composite lines for every heat sinkor source stream involved in the process-ing system, the results of heat integrationgive complex integrated systems, whichmay cause difficulty in operation and thatfinal heat integration between reactionand separation systems is carried out byengineering judgment of appropriate so-lutions through practical considerations.

Their message could be interpreted as:Keep it simple. They achieved this byadding an extra element to their synthesisstrategy the systematic addition andsubtraction of streams to and from theanalysis.

Failure to observe this important mes-sage may have cost dearly. In 1990, anengineer from a major oil company com-plained to one of us that integration wasgoing too far. Flowsheets were becomingtoo complicated. Later, he commentedthat integration had reached such a levelthat some plant operators had difficultyunderstanding where individual streamswere going.

This article demonstrates a strategybased on the keep it simple messagethat results in simpler (and cheaper) inte-gration schemes (5). With this strategy,the overall process pinch becomes lessimportant different regions of the pro-cess have their own pinch points that can

HEAT TRANSFER

CHEMICAL ENGINEERING PROGRESS DECEMBER 1999 Copyright 1999 American Institute of Chemical Engineers. All rights reserved. Copying and downloading permitted with restrictions.

Dont Let the Pinch Pinch You

Blind obedience topinch technology

rules and procedures can lead

to unnecessarily complex and

expensive, and potentially

hazardous, plants.Here we show how

plant structuresthat are both simpleand energy efficientcan be generated.

G. T. Polley and P. J. Heggs, Univ. of Manchester Institute

of Science and Technology

differ from that for the overall pro-cess. The pinch point resumes its roleas a guide to process changes ratherthan acting as the point underpinningthe design.

Problems with the pinch design method

The pinch design method has anumber of weaknesses. These can bedemonstrated using an example fromthe literature. Figure 1 is the initialnetwork derived by Ahmad andLinnhoff for a simplified aromaticsplant (6).

A recent study (7) indicates thatthe capital cost (C) of a typical singleheat exchanger (TEMA Type AEM,5-in. tube, 10-bar design pressure,etc.) is related to the total heat-trans-fer surface area (A) by the followingequation:

C = 4,600 + 920A0.7 (1)

The installed cost is about 3.5times higher than the capital cost.

This suggests a significant econo-my of scale. For instance, the fixedcost element itself would provide for

an additional 40 m2 of surface in aunit of 200 m2 nominal size. Since innetwork optimization it is often pos-sible to compensate for the removalof one unit by adding additional sur-face elsewhere, the designer shouldseek to eliminate small exchangersfrom the network in favor of largeones.

Figure 2 shows how much addi-tional surface can be purchased in an-other single exchanger with the

money saved by removing an ex-changer. The unit to be removed isrepresented by the four lines labeledA = 0 through A = 100 m2; the x-axisrefers to the size of the existing unitto which extra area is to be added,and the y-axis refers to the amount ofextra area that can be purchased.So, if a proposed 50-m2 exchanger isremoved from the design, approxi-mately 140 m2 could be added to an-other existing exchanger currently

CHEMICAL ENGINEERING PROGRESS DECEMBER 1999

HEAT TRANSFER

Source: Redrawn from (6).

40 119 220

160

60 79 189 215.8

60 103.5

17.39

159.5149.91.14

4.34

83.4327

220

220

160

300

164

138

170

300

100 201

35 141 10.1

85 3.56 3.86

18.55

85.9

1.61

60

140 6.6 141 190.3243.8

189

0.2 9.6

9.9

0.25 10.7 11.25

C 1 5 10

H

H

C

C 2 3

8 9

7

4

6

n Figure 1. Initial network designfor an aromatics plant.

300

250

200

150

100

50

025 50 100 150 200 250

Area of Existing Exchanger to Which More Surface is Added, m2

Are

a A

dded

to E

xist

ing

Exch

ange

r, m

2

300 350 400 450 500

A = Area ofExchangerRemoved

A = 100 m2

A = 50 m2

A = 25 m2

A = 0

n Figure 2. Area purchased through removal of exchangers of various sizes.

sized at 100 m2 (making the new unit240 m2), or about 180 m2 could beadded to an exchanger currently sizedat 300 m2 (making it 480 m2), beforethe purchase cost of the exchangersfor the network is increased. Thisanalysis considers only the exchangerpurchase cost; no credit has been as-sumed for cost savings associatedwith piping and other auxiliary equip-ment. Even so, a strong capital costtrade-off is indicated.

The pinch design method divides adesign optimization problem into twoparts at the pinch and produces de-signs for each subnetwork beforemerging the results to provide an ini-tial structure for subsequent optimiza-tion. By using the check-off heuristic,each subnetwork provides a designthat uses the minimum number ofunits. However, when these designsare merged, the result is an initial de-sign that uses more than the mini-mum number of units for the overallproblem.

For the aromatics plant discussedin Ref. 6, the minimum number ofunits required for the network is 10.The number in the initial design is

15. This suggests that, if installationcosts are high, this initial design is along way from the optimum configu-ration. So, lets look at how some ofthe exchangers in this design can beremoved.

Consider Exchanger 8. This is oneof the smaller exchangers in the net-work and could be a candidate for re-moval. However, it cannot be re-moved. Stream 3 must be taken to atemperature of 164C. The only otherrecovery unit on this stream is Ex-changer 3. The highest hot-streaminlet temperature possible for thisunit is 160C.

Now consider Cooler 2. This, too,is one of the smaller units. This unitcannot be removed because theneighboring heat-recovery exchangerlinks with a cold stream having a sup-ply temperature that is identical withthe required outlet temperature fromCooler 2.

This raises the question: Can anyof the coolers be removed?

The required outlet from Cooler 3is below that of the cold stream enter-ing Exchanger 2. So, this cooler can-not be removed.

The required outlet from Cooler 1is just 5 above that of the coldstream entering Exchanger 1. Notonly is this close to the practical limitfor a shell-and-tube exchanger, but itwould probably result in quite a highcapital cost penalty, making the re-moval of Cooler 1 uneconomic.

This examination indicates that theremoval of small uneconomic ex-changers can be difficult to achieveand that the initial design derivedusing the pinch design method can re-sult in nonoptimal designs.

Another weakness of the method-ology is its failure to account for thenature of the streams being processed for example, two-phase streams.Not only is the transport of two-phase streams problematic and ex-pensive (requiring large-diameterpiping), but, given the difficulty ofmaintaining good mixing of thephases, the number of heat exchang-ers in which such a stream is usedshould be minimized.

Figure 3 is the process flow dia-gram (PFD) for the aromatics plant.This diagram suggests that Streams 2and 7 are two-phase streams. In theinitial design presented in Ref. 6,Stream 2 goes through Exchangers10, 5, and 1 and through Cooler 2,and Stream 7 passes through Ex-changers 6, 3, and 2 and throughCooler 3.

Failure to relate network structureto the PFD has many implications.This becomes apparent when onecompares the proposed initial struc-ture and the PFD.

Consider Stream 4. This streamcomes from the base of first distilla-tion column and forms the feed to thesecond reactor. With this network de-sign, Stream 4 is first used to extractheat from the feed to the second gasseparator. It is then matched with theproduct from the second reactor be-fore being used to recover heat fromthe stream leaving the base of thesecond distillation column. Finally,before at last being raised to the nec-essary temperature and being fed tothe second reactor, it is matched with


DistillationColumn 1

DistillationColumn 2

Reactor1

GasSeparator

2

Reactor2

9Source: Redrawn from (6).

1 2 3

45

6

7

8

Unit E

GasSeparator

1

n Figure 3. Process flow diagram for aromatics plant.

the product from the first reactor.Now consider Stream 3, which is-

sues from the first gas separator. It isfirst matched with the product fromthe first reactor before being trans-ferred to the area of the second gasseparator, where it is matched withseparator feed. Finally, it is matchedwith the bottoms product of the sec-

ond distillation column before beingfed to the first distillation column.

Failure to recognize the impactsof network design on plant pipingand process flows is serious. It notonly significantly increases pipingcost penalties, but it also has safetyimplications.

Finally, consider the software

needs associated with the existingtechnology. Two types of programsare needed a targeting programand a network optimizer; in todayscomputer-orientated environment, itcould be argued that a network designprogram is needed as well. This is alot of software, some of which is so-phisticated. Yet, as already pointedout, the end result (despite the com-fort provided by sophisticated opti-mization) can be a nonoptimal pro-cess design.

Observations regarding network design

Networks that contain the mini-mum number of units generally havea quite simple structure that can bebroken down into a number of evensimpler self-contained subnetworks.Typical examples of such structuresare shown in Figure 4.

The structure in Figure 4a consistsof just one heat-recovery exchangerand a heater. Full use is made ofavailable process heat. There is noscope for capital cost saving, becausein terms of duty, the structure is al-ready optimal.

The structure in Figure 4b consistsof a single heat-recovery unit, aheater, and a cooler. Varying the sizeof the heat-recovery unit affects boththe capital and energy costs of thesystem, which poses an optimizationproblem. However, it is a straightfor-ward problem that can be solved byapplying integration range targetingto the two-stream problem. A sophis-ticated optimization program is notrequired.

The structure in Figure 4c hasthree process streams and two heat-recovery units plus one heater andone cooler. Here, too, an optimizationproblem exists. However, the struc-ture can be decomposed into twoparts: Exchanger A, which as a stand-alone unit does not need to be opti-mized, and a structure similar to thatalready considered (in Figure 4b),which can be optimized using a tar-geting program.

Amidpour and Polley (5) intro-


HEAT TRANSFER

First Division

Group 1Stream 1 (TAC)

Group 2Streams 2 3 4 5 6 7 8 9 (TAC)

Second Division

Group 1Streams 1 2

(TAC)

Streams 3 4 5 6 7 8 9 (TAC)

Third Division

Streams 4 5 6 7 8 9 (TAC)

Group 1 Stream 1 2 3

(TAC)

And so on...

n Figure 5. Sequential decomposition of a process integration analysis.

C

1

a.

2

C

3

b.

2

H

C

3

1

c.

2

H

n Figure 4. Three simple subnetworks.

duced problem decomposition intoprocess integration analysis. In thisprocedure, an overall integrationproblem is decomposed into a num-ber of self-contained zones in accor-dance with the PFD. Integration isthen restricted to streams withinzones unless integration betweenzones was clearly economic. In an-other paper (8), they showed howpiping constraints could be dealt withby ensuring that the more-difficultstreams (such as two-phase streams,condensing vapors, evaporating liq-uids, and low-pressure gases) are nottransferred out of their zones.

By decomposing the overall prob-lem in such a manner, the engineerends up designing networks that arelocal in terms of operation, whichtherefore can easily be understood byoperators. They are also associatedwith local unit operations and thusare unlikely to result in unnecessarilyexpensive pipe runs. Finally, ratherthan dealing with a single large net-

work-design problem, the engineer ishandling a small number of subprob-lems dealing with fewer streams. Thedesign of these smaller individualnetworks is likely to be a simpler taskthan developing a network for thewhole process. The designer willoften be able to solve the series ofsubproblems faster, and to better ef-fect, than the overall problem.

Problem decomposition procedure

The following procedure for prob-lem decomposition (Figure 5) can beapplied using simple process-integra-tion targeting software:

1. Number the streams sequentially.2. Apply range targeting to the full

stream set and establish a referencetotal annual cost (TAC).

3. Split the stream set into twoparts, the first set consisting ofStream 1 and the second set consist-ing of the remaining streams (here,Streams 2 to 9).

4. Apply range targeting to bothsets and sum the results.

5. Derive new sets by changingthe position of the boundary, so thatthe new first set contains Streams 1and 2 and the new second set con-tains the remaining streams (Streams3 to 9).

6. Repeat Steps 4 and 5 until all ofthe boundary positions have been examined.

Stagewise application of problem decomposition

The order in which the streams arelisted affects the results of the decom-position analysis. By arranging themin the order in which they appear onthe PFD, the designer can identifycost-effective local integration (as op-posed to cross-plant integration). Thishas benefits in terms of piping cost,plant operability, and plant simplicity.But, once these opportunities are ex-hausted, this ordering has no benefit.

For general thermal integration,


4

6

3

1

2 5 9 13

12

108

11

15

16 20

17

19

18

22

21

14

T2 T3T1 T6 T4

R1

T5 R2 T10 T7

n Figure 6. Process flow diagramfor the aromaticsplant in the example.

the streams should be ordered in linewith the Ponton-Donaldson Heuristicof: Match the hot stream havingthe highest supply temperature withthe cold stream having the highesttarget temperature. (9)

In applying this approach, therewill be occasions where the residu-al of the hot stream (i.e., the hotstream condition after a match hasbeen made) may be important. So, theanalysis is conducted in the followingstages:

1. Order the streams in line withthe PFD and identify the local inte-gration opportunities.

2. Having exhausted the opportu-nities for cost-effective local integra-tion, reorder the remaining streams inline with the Ponton-DonaldsonHeuristic and repeat the analysis.

3. If Step 2 results in significantpotential cost penalties, identify thehot stream residuals that influence theheat recovery and subdivide these hotstreams (still ordering in line with thePonton-Donaldson Heuristic).Overall ordering of a process integration study

A process integration analysisusing this approach involves the fol-lowing five stages.

1. Problem simplification re-moval of streams from problem.

2. Identification of processchanges improved overall efficien-cy through energy and utility costsavings, and capital cost reduction.

3. Setting of final problem se-lection of utility levels, specificationof utility costs, selection of heat-transfer parameters, and adjustmentof stream ordering, if desirable.

4. Decomposition analysis on thebasis of the layout identification ofcost-effective local integration.

5. Decomposition analysis on athermal basis development of theremainder of the network.

Example: aromatics solvent plant

Figure 6 is the PFD for an aromat-ics solvent plant. We will use it to

demonstrate the basic analysis tech-nique. Consequently, we will notcover some of the more complex is-sues such as pressure drop specifica-tion, consideration of exchanger tech-nology, and so on. The problem istherefore defined in terms of streamheat capacity flow rates, and physicalproperty information is not required.The streams have been listed in theorder in which they appear on thePFD. The stream data (heat capacityflow rate (CP), supply temperature(Ts), target temperature (Tt), and load)are summarized in Table 1.

Stage 1 Simplify the problem

The stream data are entered into aprocess-integration targeting pro-gram. (We have used INTEGRITY,developed by ESDU, which alreadyincorporates the problem decomposi-tion procedure.)

The hot utility temperature is set at500C, and the cold utility tempera-ture is set at 20C. A point target, at aminimum temperature approach of10C, is generated. Examination ofthe HDS diagram (Figure 7) showsthe following:

one overhead condenser is belowbut close to the pinch (examining thetemperatures and loads in Table 1,this is identified as overheads on T2);

two reboilers are above but quiteclose to the pinch (again using thedata in Table 1, these are identified asthe reboilers on T6 and T10);

the remaining reboilers are wellabove the pinch and, given the shapeof the HDS diagram, which shows nohigh-temperature heat surplus, thesecannot be driven through heat recov-ery; and

the remaining condensers arewell below the pinch and, given thatthe HDS diagram indicates no low-


HEAT TRANSFER

Stream Name CP, kW/C Ts, C Tt, C Load, kW

1 Make-up 8 20 80 4802 T1 Tops 72 38 80 3,0243 T1 Bottoms 30 38 49 3304 T2 Reboiler 12,000 140 141 12,0005 T2 Condenser 12,400 66 65 +12,4006 T2 Bottoms 14 140 30 +1,5407 T3 Feed 46 65 38 +1,2428 T6 Reboiler 7,600 79 80 7,6009 T6 Condenser 5,900 56 55 +5,900

10 Raffinate 46 80 30 +2,30011 T4 Preheat 16 38 80 67212 T4 Reboiler 4,100 120 121 4,10013 T4 Condenser 3,500 59 58 +3,50014 T4 Bottoms 18 121 38 +1,49415 T5 Reboiler 8,000 110 111 8,00016 T5 Condenser 9,500 59 58 +9,50017 R2 Feed 75.2 88 204 7,56418 R2 Effluent 52.8 188 8 +5,22819 T10 Reboiler 2,200 66 67 2,20020 T10 Condenser 2,000 48 47 +2,00021 T7 Reboiler 3,800 114 115 3,80022 T7 Condenser 3,460 50 49 +3,46

Table 1. Stream data for the example.

temperature deficits, heat recoveryfrom these units is not justified.

Therefore, assume that the reboil-ers on Columns T2, T4, T5, and T7are driven by hot utility, and deletethese streams from the problem. Alsoassume that the condensers onColumns T4, T5, T6, T7, and T10 arecooled using cold utility, and deletethese streams from the problem.

Stage 2 Identify process changes

The condenser on Column T2 isbelow and close to the pinch whilethe reboilers on Columns T6 and T10are above and close to the pinch. Inaddition, the load on the condenser isclose to the combined reboiler loads.

The pressure at which Column T2operates can be increased. This wouldresult in an increase in the tempera-ture at which the overheads condense.It may be possible to use this vapor todrive the two reboilers. If the temper-ature of the overheads could be raisedhigh enough, very-low-pressure

steam could be generated and thisused to drive the reboilers. An alter-native approach would be to reducethe pressures on Columns T6 andT10. Evaluation of these alternativesrequires the use of a process simula-tor and is outside the scope of this article.

If process modification is not vi-able, the HDS diagram indicates thatthe condenser should be cooled usingcold utility and the reboilers drivenusing hot utility.

The following actions are there-fore recommended:

1. Examine possible process modi-fications. (Since the alternative is torun both units using utilities, this canbe done after the rest of the heat-re-covery system has been designed.)

2. Delete all of the condensers andreboilers from the problem.

Stage 3 Set the final problem

Utility levels and costs. Examina-tion of the HDS diagram suggests

that the best form of hot utility issteam. Heat must be provided up to atemperature of 204C. This suggestsa steam temperature of around220C.

The hot utility cost is then calcu-lated as follows. Cost of fuel oil =$0.16/L; potential heat extractionfrom fuel = 40.8 MJ/L; base cost (in-cluding 10% to account for capitalcosts and assuming that the fuel isused at 85% efficiency) =1.294(0.16)/40.8 = $0.0051/MJ. Tem-perature factor (which accounts forthe variation in the heat of vaporiza-tion of steam with temperature) =1.043 + (2.407 10-4)(220) (5.26 10-6)(220)2 = 0.844. Steam cost =0.0051/0.844 = $0.00604/MJ =$0.0218/kWh.

Examination of the HDS diagramindicates that cooling water (at 20C)is a suitable cold utility. The coolingwater cost is assumed to be$0.0016/kWh.

Capital costs assumptions. Thecapital cost of an exchanger is givenby:

C = a + bAc (2)

The values of the constants a, b,and c for a typical carbon steel ex-changer are given in (7). The valuesused here are a = 16,000, b = 3,200,and c = 0.7.

Annualization factors. The fol-lowing annualization factors are as-sumed: 8,000 h/yr operation, 10-yrplant life, and 10% interest rate.

Heat exchanger parameters.Multipass exchangers are assumed,and the X factor (that is, the limitplaced on the approach to maximumthermal effectiveness (10)) is set at0.9. The maximum exchanger size isset at 500 m2/shell.

Typical stream heat-transfer coef-ficients and fouling resistances aredetermined (for example, using Ref.11). For liquid organic (light)streams, Ref. 7 suggests approximatefilm heat-transfer coefficients of1,500 W/m2K and a fouling resis-tance of 0.0004 m2K /W. Combining


300

350

250

200

150

T10

T10

T6

T5

T5

T7

T7

T4

T4

T2

T2

100

50

05,000

Minimum temperature difference = 10C

Inte

rnal

Tem

pera

ture

, C

10,000 15,000 20,000 25,000 30,000 35,0000

Enthalpy, kW

n Figure 7. Heat demand and supply diagram for the aromatics plant in the example.

the two resistances and then invertingyields a duty coefficient of 937.5W/ m2K.

For steam, a coefficient of 7,000W/ m2K is recommended with afouling resistance of 0.0001 m2K /W.Combining gives a duty coefficient of4,120 W/ m2K, so a value of 4,000W m2K will be used. For coolingwater, a coefficient of 4,000 W/ m2Kis suggested with a fouling resistanceof 0.0002 m2K /W. Combining givesa duty coefficient of 2,220 W/m2Km2K, so a value of 2,000 W/ m2Kwill be used.

Stage 4 Decompositionbased on layout

The streams are ordered in accor-dance with their appearance on thePFD. A range target for the full prob-lem is then determined. The mini-mum TAC is found to be$581,000/yr. This is a theoreticalminimum cost. The TAC of the finaldesign can be expected to be 1015%higher than this.

Range targeting can now be ap-plied systematically to different datasets to identify how the problem canbe cost-effectively decomposed intolocal integration problems.

Decomposition into two parts, onegroup containing Streams 1 to 8 andone group containing Streams 9 and10, yields a system with a theoreticalminimum TAC of $611,000/yr. Thisis made up of a theoretical cost of$194,000/yr (at an optimum approachtemperature difference of 15C) forthe first group and an actual opti-mized cost of $417,000/yr (at an opti-mum temperature difference of 10C)for the heat-recovery system involv-ing the reactor streams (Streams 9and 10).

Comparing this decomposed costof $611,000/yr with the theoreticalminimum of $581,000/yr indicates apotential cost penalty associated withthis real match of $30,000/yr. This is5.2 % of the theoretical minimumcost.

Separate the reactor feed stream(Stream 9) and reactor effluent stream


HEAT TRANSFER

Local integration around reactorTAC = $147,000/yrPotential Penalty = $3,000/yrHeat-recovery match (reactor feed against reactor effluent) requires 714 m2

Heater on reactor feed area = 114 m2

Cooler on reactor effluent area = 10 m2

Local integration around column T4TAC = $25,000/yrPotential Penalty = $35,000/yrHeat-recovery match (T4 bottoms against T4 preheat) requires 39 m2

Cooler on T4 bottoms area = 36 m2

Make-up heaterTAC = $59,000/yrPotential Penalty = $2,000/yrArea = 4 m2

Heat recovery between T3 feed and T1 bottomsTAC = $28,000/yrPotential Penalty = $12,000/yrHeat-recovery exchanger area = 41 m2

Cooler on T3 feed area = 55 m2

Heat recovery between raffinate and T1 topsTAC = $55,000/yrPotential Penalty = $16,000/yrHeat-recovery exchanger area = 241 m2

Cooler on raffinate area = 77 m2

Heat recovery between T2 bottoms and T1 topsTAC = $92,000/yrHeat-recovery exchanger area = 75 m2

Heater on T1 tops area = 5 m2

Cooler on T2 bottoms area = 36 m2

Final TAC = $676,000/yr

Table 3. Summary of final design for the example.

Stream Ts, C Tt, C Load, kW

T1 Tops 38 80 3,024T1 Bottoms 38 49 330 T2 Bottoms 140 30 1,540T3 Feed 65 38 1,242Raffinate 80 30 2,300

Table 2. Remaining five streams to undergo decomposition on a thermal basis.

(Stream 10) from the problem and de-velop a separate subnetwork for thissystem. It consists of a heater having114 m2 of surface, a heat-recoveryexchanger of 714 m2, and a cooler of10 m2.

Decomposition analysis of the re-maining eight streams is now under-taken. The minimum TAC for thesubsystem is $194,000/yr. The analy-sis indicates that the problem canagain be broken down into two sec-tions, the first containing Streams 1 to6 (TAC = $204,000/yr) and the sec-ond containing just Streams 7 and 8(TAC = $25,000/yr). Here, too, thesecond group is an actual heat-recov-ery match. The TAC for the decom-posed problem is $229,000/yr, with apotential cost penalty of $35,000/yr(6% of the theoretical minimum).

Separate the T4 preheat (Stream 7)and the T4 bottoms (Stream 8) fromthe problem and integrate them sepa-rately. The optimum temperature dif-ference for this subsystem is 41Cand the associated TAC is $25,000/yr.The network consists of two units, aheat-recovery unit of 39 m2 and acooler of 36 m2.

Decomposition analysis of the re-maining six streams is now undertak-en. First consider using utility to heatthe make-up stream (Stream 1). TheTAC for the six-stream problem is, asnoted above, $204,000/yr. For the de-composed problem, the TAC is$206,000/yr.

Heat the make-up stream using hotutility. This requires a heater of 4 m2and has a TAC of $59,000/yr.

Decomposition analysis is appliedto the remaining five streams. Allpossible decompostions are found toinvolve significant penalty. Thus, at-tention should now be directed to ap-plying decomposition analysis on athermal basis.

Stage 5 Decomposition on a thermal basis

The five streams remaining arelisted in Table 2. These streams arenow ordered in accordance with thehottest target temperatures for the

cold streams and hottest supply tem-peratures for the hot streams. Tem-perature spans and stream heat loadsprovide secondary guidance with re-gard to the ordering.

Stream 1 should be T1 tops (thecold stream with the hottest targettemperature). The next stream shouldbe the hot stream with the highestsupply temperature, which is T2 bot-toms. This hot stream has insufficientload to fully handle the cold stream.So, the next stream should also be ahot stream. Of the remaining hotstreams, the raffinate has the highestsupply temperature. The remainingtwo streams can be placed in eitherorder. We choose to place the hotstream first. So, the ranking is now:

Stream 1 = T1 tops; Stream 2 = T2 bottoms; Stream 3 = raffinate; Stream 4 = T3 feed; and Stream 5 = T1 bottoms.Decomposition analysis is then ap-

plied to this order. The TAC of theoverall subsystem remains at$147,000/yr. However, there is now asuitable decomposition involving asystem containing Streams 1 to 3 andone containing an actual match be-tween Streams 4 and 5. The actualcost of the match between the T3 feedand the T1 bottoms is found to be$28,000/yr. The minimum theoreticalcost for the three-stream system is$131,000/yr. The combined costs are$159,000/yr, which has a potentialcost penalty of $12,000/yr.

Match Streams T3 feed and T1bottoms. A point target exists at theidentified optimum approach of 16C.

The network contains two units, aheat-recovery exchanger of 41 m2 anda cooler of 55 m2.

Now just three streams are left.The optimum temperature approachfor the subnetwork involving thesestreams is 10C and the TAC is$131,000/yr. A point target at thistemperature approach indicates thatthe pinch occurs at a cold stream tem-perature of 38C and that 264 kW ofhot utility and 1,080 kW of cold utili-ty are required. Application of thepinch design method to this problemwould result in a network containinga stream split as shown in Figure 8.

An alternative serial arrangement isshown in Figure 9. First, the raffinate(the hot stream with the supply tem-perature closest to the pinch) ismatched against T1 tops. Range tar-geting indicates an optimum tempera-ture approach of 10C. The area re-quired for the raffinate cooler is 77 m2,and the area of the heat-recovery ex-changer contacting the raffinate withT1 tops is 241 m2. The remaining dutyon the T1 tops stream is 1,552 kW.Thus, the temperature of the streamleaving the heat-recovery unit is: Tt


n Figure 9. Alternative serial structure.

Raffinate

T2 Bottoms

T1 Tops

1288

828

147238 80

140

80

4830

4830

264

252

n Figure 8. Structure derivedusing the pinchdesign method.

1,552/CP = 80 1,552/72 = 58.4C.The T1 tops stream is now divided

into two parts: the first element in-volves heating the stream from 38Cto 58.4C, the second heating from58.4C to the final target temperatureof 80C. Matching the raffinate withthe first element results in a subsys-tem of with a TAC of $55,000/yr. Op-timizing the match between the re-maining part of the T1 tops streamand the T2 bottoms stream indicatesan optimum temperature approach of10C and a TAC of $92,000/yr. TheT2 tops cooler has an area of 36 m2,the remaining heater has an area of 5m2, and the heat-recovery exchangerhas an area of 75 m2.

The overall TAC of the serial ar-rangement is $147,000/yr. This isclose to the minimum value of$131,000/yr, which would require thestream-split arrangement. The serialdesign is accepted.

The final design is summarized inTable 3. The final TAC is $676,000/yr,which is 16.4% higher than the theo-retical minimum cost. The differencesare accounted for as follows: integra-tion around the reactor: 5.2%; integra-tion around T4: 6.0%; make-upheater: 0.3%; T3/T1 recovery: 2.1%;the serial subnetwork: 2.8%; and theapplication of local integration:11.5%. The piping costs associatedwith the remainder of the schemehave not been determined.

Implications for process integration software

The procedures described herehave major benefits. They allow thedesigner to produce energy-efficientnetwork structures that are simple.They control piping costs and avoidthe hazards associated with over-inte-gration. They simplify the networkdesign process. They avoid the needfor sophisticated network-optimiza-tion procedures and can be employedusing any process-integration target-ing program. This has implicationsfor software vendors and purchasers.

Currently, process integration soft-ware falls into two categories. There

are expensive programs that providetargeting, network design, networkanalysis, and network optimization.

There are much-less-expensive pro-grams that concentrate on targetingcalculations. The techniques present-ed here favor these less-expensiveprograms. [See the CEP Online Soft-ware Directory at http://www.aiche.org/software/softwareindex.htm forinformation on specfic software pack-ages. Editor] CEP


HEAT TRANSFER

Literature Cited1. Umeda, T., et al., Heat Exchange Sys-

tem Synthesis, Chem. Eng. Progress, 74(7), pp. 7076 (July 1978).

2. Umeda, T., T. Harada, and K. Shiroko,A Thermodynamic Approach to theSynthesis of Heat Integration Systems inChemical Processes, Computers andChemical Engineering, 3, pp. 273282(1979).

3. Umeda, T., K. Niida, and K. Shiroko,A Thermodynamic Approach to to HeatIntegration in Distillation Systems,AIChE Journal, 25 (3), pp. 423429(May 1979).

4. Itoh, J., K. Shiroko K., and T. Umeda,Extensive Use of the T-Q Diagram toHeat Integrated System Synthesis, pre-sented at the International Symposiumon Process Systems Engineering, Kyoto,1982, and published in Computers andChemical Engineering, 10, pp. 5966(1986).

5. Amidpour, M., and G. T. Polley, Appli-cation of Problem Decomposition inProcess Integration, Transactions ofIChemE, 75A, pp. 5363 (1997).

6. Ahmad, S., and B. Linnhoff, Supertar-geting: Different Process Structures forDifferent Economics, Journal of Ener-gy Resources Technology, 111, pp.131136 (1989).

7. ESDU International, PLC, Costing ofShell-and-Tube Heat Exchangers,ESDU, London (1999).

8. Amidpour, M., and G. T. Polley, Deal-ing With Piping Constraints in Heat Ex-changer Network Synthesis, IChemESymposium on Process Integration andFluid Separations, Manchester, U.K.(June 1994).

9. Ponton J. W., and R. A. B. Donaldson,A Fast Method for the Synthesis of Op-timal Heat Exchanger Networks, Chem.Eng. Sci., 29, pp. 23752377 (1974).

10. Ahmad S., B. Linnhoff, and R. Smith,Design of Multipass Heat Exchangers:An Alternative Approach, Journal ofHeat Transfer, 110, pp 304309 (May1990).

11. Saunders, E. A. D., Heat Exchangers:Selection, Design and Construction,Longman Group, London, copublishedin the U.S. by John Wiley & Sons, NewYork (1988).

G. T. POLLEY (Phone: +44-1229-585-330; Fax: +44-1229-585-708; E-mail:[email protected]) is now retiredfrom the Univ. of Manchester Institute ofScience and Technology, Manchester, U.K.,and is consolidating and publishing hisresearch. Following PhD research intocondensation heat transfer, he joined theU.K.s National Engineering Laboratory in1975, and was subsequently responsible forresearch into boiling-heat-transfer and heat-recovery system design. Since 1978, he hasbeen a member of the International EnergyAgency Working Party on Heat ExchangerTechnology. In 1985, joined the chemicalengineering department at the Univ. ofManchester Institute of Science andTechnology, as the director of the Centre forProcess Integration. His own researchactivities centered around the development ofprocess and equipment designmethodologies. In 1990 was awarded theMoulton Medal by the Institution of ChemicalEngineers in recognition of work on oilrefinery revamping, and in 1992 was awardedthe Ackrill Trophy by the U.K. Heat TransferSociety for work on the applications of heat-transfer enhancement. He holds BTech, MSc,and PhD degrees in chemical engineeringfrom Loughborough Univ. of Technology, andis a member of AIChE.

P. J. HEGGS is a professor and head of thechemical engineering department at the Univ.of Manchester Institute of Science andTechnology, Manchester, U.K. (Phone: +44-0161-200-4370; Fax: +44-0161-200-4399; E-mail:[email protected]). Previously, he heldpositions as a professor at the Univ. ofBradford, a lecturer at Leeds Univ., and asenior engineer at Union Carbide Corp.sTechnical Center in South Charleston, WV. Hehas been involved in heat-transfer researchfor over 35 years, and his other interestsinclude adsorption/desorption, reactionengineering, and mathematical modeling ofengineering equipment and processes. Heholds a BSc and PhD in chemical engineeringfrom Leeds Univ., and he is a Fellow of theInstitution of Chemical Engineers, a Fellow ofthe Royal Academy of Engineering, and aChartered Engineer.

don't let the pinch

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