Screenging of Process Retrofit Alternatives

8/2/2019 Screenging of Process Retrofit Alternatives

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I n d . Eng. C h e m . R e s . 1987,26, 2195-2204 21951 0

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0 2 4 6 8 1 0 0 4 8 12 16 20se t Se tFigure 9. Circulation time distributions of pill and liquid liquidvolume, 19 m3; tw o impellers; gas rate, 2.6 cm/s; detection withaerials in compartments 2 and 4.partments 2 + 4 has been calculated.Finally the circulation time distribution of the liquid wascalculated in the same way by setting the rate of fall tozero. The observed and the calculated distributions of thepill aswell as he calculated distributions of the liquid aregiven in Figures 3-9.4. DiscussionThe simple five-compartments model gives a good fitfor the tail of the circulation time distribution of the pill.In all cases, the fit for the short circulation times is notas good.There is a time lag in the beginning of the observed pillcirculation time distribution curves, which is not explainedby the model. Apparently the assumption of the liquidphase in each compartment being ideally mixed is toosimple. The pill, having left a detection compartment, hasat least a minimum residence time in the next compart-ment before reentering the same aerial compartment. Inall experiments, the observed pill circulation time dis tri-bution shows one or more peaks; this will also be causedby the nonideality of the mixing in the various compart-ments.In ou r model, we assume that the probability for the pillto leave a compartment by falling out or by following theliquid flow is independent of the path the pill has beenfollowing before entering the compartment. For a noni-deally mixed system, this will not be completely true.The shape of the observed distribution curves is thesame as the shape of the curves reported by Mann et al.(1981).After including two impellers and a rate of fall for thepill, their structured stochastic flow model may be useful

for interpreting our measurements. Such a model containsa large number of ideal mixers in series and parallel andis likely to be capable of fi tting both short and long cir-culation times.After some modifications, the presented model can beused for the calculation of gas holdup, residence timedistribution in solid-liquid dispersions, etc.Acknowledgment

The mechanical construction of the radio pill and theaerials was designed and developed by H. van Dam. Theelectronics were designed and built by L. P. de Meul-meester, who made a major contribution to the measure-ments.Nomenclaturefi+j = frequency with which the pill, being in compartmentf , = frequency for a pill leaving a compartment with liquidf, = frequency for a pill leaving a compartment by falling out,Hi height of compartment i, mp , = falling probability of a pillq = probability that the pill is present in compartment iQCJ= liquid circulation capacity of an impeller, m3/st p time, stc f = time between two successive passages of a liquid elementtc,: = time between two successive passages of the pill throughv , = superficial gas velocity, m/sv , = rate of fall of a pill within a compartment, m/sV = compartment volume, m3G re e k Symbol st ,= volumetric liquid fractionTI , = mean residence time of the pill in compartment i, sT{ = mean residence time of a liquid element in compartmentT: = mean time for the pill to fall out of a compartment, sLiterature CitedBarneveld, J. v.; Oosterhuis, N. M. G .;Pragt, J. J.; Smit, W. Znd.Mann, R.; Mavros, P. P.; Middleton, J. C. Tran s. Znst. Chem. Eng.

Received for review April 21, 1986Revised manuscript received May 19, 1987Accepted June 11, 1987

i, moves toward compartment j , s-lflow, s-1S-1

through compartment i, scompartment i, s

i, s

Eng. Chem. Res. 1987, preceding paper in this issue.1981,59, 271-278.

Screening of Process Retrofit AlternativesWayne R. Fisher, Michael F. Doherty, and James M. Douglas*Department of Chemical Engineering, Universi ty of Massachuse t ts , Amherst , Massachuse t ts 01003A systematic procedure for developing and screening process retrofit opportunities is presented. Th eprocedure considers both modifications in the structure of th e flow sheet and in equipment sizesfor a fixed flow sheet. Also, a systematic way of identifying bottlenecking equipment is described.Th e results of case studies indicate that retrofitting to reduce raw materials costs often is much moreimpor tant tha n retrofitting to save energy.

Expert systems and intelligent computer codes haveth e potential of making a dramatic impact on the practiceof chemical engineering. As a minimum, we expect tha tsolution strategies for various types of problems will becombined with hierarchies of algorithms to make it possible0888-5885/87/2626-2195$01.50/0

to explore more process alternatives quickly and inex-pensively. A t present, knowledge engineers attempt tocapture solution strategies by interviewing experts andwatching them work. However, it is possible to develophierarchical, top-down, least-commitment strategies0 1987 American Chemical Society


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2196 Ind. Eng. Chem. Res., Vol. 26, No . 11,1987th at provide efficient solutions in a variety of other ways.Westerberg (1986)recently reviewed four strategies of thistype that have been proposed for synthesizing process flowsheets. In this paper, we describe a solution strategy forretrofit problems.There has been rapidly growing interest in process re-trofits in recent years. By retrofitt ing we mean makingminor changes in the structure of a flow sheet and/or someequipment sizes in order t o (a) significantly reduce oper-ating costs, (b) increase the capacity of the process, (c)process new feedstocks, and /or (d) incorporate a newtechnology (i.e., a new catalyst, a new membrane separator,etc.). In part, this recent interest may be attributed to theimpact of changing economic forces on the chemical in-dustry, and in part it may be attributed to the great successof applying the new heat integration technology to retrofitsituations. That is, Imperial Chemical Industries (Bolandand Hindmarsh, 1984) and Union Carbide (Linnhoff andVredeveld, 1984) have shown that often 30-5070 savingsin energy are possible even in retrofit applications. Otherstudies indicate similar trends (Steinmetz and Chaney,1985; Witherell and Linnhoff, 1985). Hence, an incentivefor retrofitting has been clearly established.The very successful energy integration studies discussedabove have been based on the process flows th at exist inthe plant. However, the design variables that fix theprocess flows (e.g., conversion, molar ratio of reactants,purge composition, etc.) normally involve economic tra-deoffs which balance incremental raw material costs dueto selectivity losses of reactants to byproducts againstincremental recycle costs. Of course, the recycle costs aredependent on the heat-exchanger network so that theproblems of finding the optimum process flows andheat-exchanger network are coupled.The coupling of the optimum process flow problem withthe optimum energy integration implies that retrofitstudies which only consider energy may not give the bestretrof it design. For example, if an energy retrof it studyindicates that large energy savings are possible and thatthese energy savings have a significant effect on the recyclecosts, then it should be possible to convert some, or pos-sibly all, of the energy savings into raw material savingsby increasing the recycle flows. For a large number ofprocesses, raw material losses to byproducts are much moreimportant than energy costs, and therefore there is greatincentive for developing a strategy of process retrofits thatconsiders process flows, as well as energy.Goals of This Research

Since a retrofit analysis requires looking at the completeplant, it is a large and complicated problem. Thus, in orderto minimize the engineering effort required to identify themost important retrofit opportunities, it would be desirableto have approximate methods available that could be usedfor screening purposes. Moreover, it would be highly de-sirable to develop a general, systematic procedure basedon these shortcut techniques, which could provide the basisfor an interactive computer code that could be used toscreen retrofit opportunities. Our goal in this research wasto establish a systematic procedure of this type. Fortu-nately, many of the tools th at we have developed for thepreliminary design of new processes are also useful forretrofit studies.A Systematic Procedure for Retrofitting

The retrofit procedure that we describe below has beendeveloped in terms of a hierarchical set of problems. Thishierarchy corresponds to establishing a series of bounds

Table I. Retrofit Screening Strategy1. Use an operating cost diagram t o identify the incentive for rawmaterials and energy savings.2. Determine the incentive for completely replacing the existingplant.a. Estimat e the op timu m values of the design variables withcurre nt costs.b. Identify important process alternatives.

we completely replace the plant.or a structural alternative.a. Eliminate the existing heat exchangers, but retain theheating and cooling ultilities costs.b. Identify the dominant operating variables.c . Identify th e equipment th at constrains the dominantoperating variables.d. Remove the equipment con straints by adding incrementalequipment capacity until the incremental investment costsbalance the increm ental savings in operating costs.e. Develop a new heat-exchanger network for the process.f. Modify the new heat-exchanger network in order to use asmuch existing heat-exchange equipm ent as possible.g. Reoptimize the process flows and heat-exchanger network.

3. Screen the process alternatives, and fin d the best flow sheet if4. Modify the existing equi pme nt sizes for the existing flow sheet

5 . Refine the retrofit calculations.and/or targets, which make it possible to terminate theretrofit study early if there is not a sufficient economicincentive to justify additional effort. Thus, the hierarchycorresponds to a top-down, least commitment strategy,and consists of the five levels given in Table I. Each stepin the hierarchy is discussed in more detail below.1. Identify the Incentive for Raw Material andEnergy Retrofitting. The first step in our procedure isto prepare an operating cost diagram. That is, all thesignificant operating costs are attached to stream arrowson the current process flow sheet. In particular, we showthe following.a. We show the cost of steam or fuel for all reboilers,furnaces, preheaters, etc.b. We show the cost of electric power (or steam) for allcompressors and blowers. Normally, we neglect the op-erating costs of pumps if they are less than 10% of themost expensive operating cost; Le., we only want to con-sider the most important operating costs in order to sim-plify the screening calculations.c. We show any other significant energy-related oper-ating costs.d. We show the cost of raw materials in excess of thestoichiometric requirements to achieve the existing pro-duction rate (we use the maximum value if the productionrate is changing). For complex reactions, the raw materialrequirement will depend on the selectivity losses to by-products and the purge losses that exist in the currentplant.e. We show the credit associated with any byproductstreams that are sold or are used as fuel and the pollutiontreatment cost of all waste streams.This operating cost diagram is based on the cost dia-grams that Douglas and Woodcock (1985) describe forprocess synthesis and the quick screening of process al-ternatives. From an inspection of this operating cost di-agram, we can quickly evaluate the maximum incentivefor making process modifications in order to reduce theraw material losses to byproducts, the energy costs, or both.For complex reactions, the raw material losses are usu-ally dominant. The relative value of the coproducts dic-tates the desired operation of the process. In some in-stances this corresponds to a single valuable productbalanced against byproducts which have either waste orfuel value. In other cases, there will be several products(e.g., glycols), each of differing value. However, if there


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Ind. Eng. Chem. Res., Vol. 26, No . 11, 1987 2197is only small potential for raw material savings, but a largepotential for energy savings, then we bypass the remainderof this procedure and just focus on energy retrofitting byusing the new energy integration procedures that havebeen published recently, i.e., Linnhoff et al. (1982), Bolandand Hindmarsh (1984), Linnhoff and Vredeveld (1984),Tjoe and Linnhoff (1986).

Determine the Economics of Replacing theWhole Plant for the Existing Flow Sheet. One optionwe have is to replace the complete plant. We want toevaluate this option (assuming we can do it quickly) bothto provide a benchmark for our ret rofit analysis and alsoto evaluate the costs and profitability position of ourcompetitors if they built a process like ours. In the nextstep, we will also consider the cost of building a new plantthat corresponds to the best process alternative.We can use the hierarchical decision procedure describedby Douglas (1985) to estimate the optimum design of a newplant. This procedure usas approximate material andenergy balances, as well as shortcut equipment and costmodels,so that an estimate of the optimum design can beobtained, within 2 days to 1week. An interactive computerprogram developed by Kirkwood et al. (1987) that is basedon this procedure makes it possible to complete a con-ceptual design in about 1 h.Since the hierarchical design procedure is systematic,it makes it a simple matter to generate a list of processalternatives. These are the same alternatives that we wantto consider for structural modifications of the process flowsheet in our retrofit analysis. Thus, using the hierarchicaldesign procedure to redesign the existing plant has fourpositive contributions: (a) it helps the user to quicklyunderstand the process, (b) it provides a current pictureof t he design economics, (c) it generates a list of processalternatives, and (d) it estimates the optimum values ofthe most significant design variables with current costs.Of course, if the new values of the optimum design vari-ables and costs are close to the existing values, then thereis little incentive for continuing to consider the retrofit ofthe process.3. Screen the Process Alternatives for a New De-sign. A procedure for the quick screening of process al-ternatives for a new design has been presented by Douglasand Woodcock (1985). Star ting with a base-case (opti-mum) design, the procedure allocates the costs of heatexchangers o the process streams based on the reciprocalsof the heat-transfer coefficients of each stream (Townsendand Linnhoff, 1984). Then, costs are allocated to the freshfeed streams, gas recycle streams, and liquid recyclestreams based on the process flows and heat loads. Next,the effect that changing the structure of the flow sheet hason each of the process flows is estimated, and the effectsof these flow rate changes on the fresh feed, gas recycle,and liquid recycle costs are calculated. For some alter-natives, e.g., an equipment replacement, the process flowsare not changed, but a comparison between alternativesis straightforward.The goal of this rough screening analysis is to estimatethe order of magnitude of savings obtained by using var-ious process alternatives without having to redesign theprocess. In other words, we attempt to find the incentiveto go through the hierarchical design procedure again forthe various alternatives. We repeat the hierarchical designprocedure for the most promising alternatives in order toestimate th e optimum design conditions for each alter-native. By comparing the optimum design for a new plantby using our existing flow sheet with the optimum designof the best alternatives, we generate a list of promising

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x (conversion)Figure 1. Typical cost tradeoffs for each design variable in a flowsheet.modifications for the existing flow sheet that should beconsidered in a retrofit analysis.Of course, if a retrofit alternative does not use an existingpiece of equipment in the plant, we must still include thecontinuing payment for that equipment in our retrofiteconomics. Furthermore, if our most promising alternativerequires changing the process flow rates, we might en-counter equipment constraints in our existing plant. Inorder to better understand the nature of these equipmentconstraints, it seems desirable to deviate from our retro-fitting procedure to review the difference between de-signing to achieve the minimum total annual cost (i.e., thesum of annualized capital and operating cost) vs. at-tempting to retrofit a design to achieve the minimum op-erating cost.Optimum Design vs. Minimizing Operating Costs.Most design problems have some processing constraints.For example, if the reactor temperature exceeds a certainvalue, undesirable side reactions may take place. Similarly,if the molar ratio of reactants is less than a certain value,undesirable side reactions, such as coking, may occur. Ifwe were willing to undertake more detailed kinetic mod-eling, we could treat these situations as optimum designproblems, but normally it is much simpler to treat themas constraints. We can satisfy these constraints by re-moving one of the optimum design variables (degrees offreedom) from the design problem (by using variableelimination, fo r example).For each of the remaining design variables, we normallyfind that there is a minimum (annualized) capital cost, aminimum operating cost, and a minimum total annualizedcost (annualized capital plus operating cost); see Figure1. We do not design to obtain the minimum operatingcost because the capital investment is too high, nor do wedesign to obtain the minimum capital cost because theoperating costs are too high. Thus, we trade off annualizedcapital vs. operating cost, and we design for the minimumtotal annual cost. Of course, in many cases this optimumdesign is not operable when disturbances enter the plant,and in these situations we must include some overdesign.This problem of optimal overdesign to accommodatedisturbances is discussed elsewhere (Fisher et al., 1987).As soon as the plant is built, we have spent ou r totalcapital investment. From that time on, we would like toobtain our desired production rate and product purity, aswell as satisfy the process constraints, with the minimumoperating costs. Thus, i t is always desirable to change theoperating variables away from their optimum steady-statedesign settings in the direction that minimizes operatingcosts; see Figure 1. When we change the values of theseoperating variables, we normally encounter equipmentconstrainsts because the equipment sizes were fixed by thebest capital vs. operating cost tradeoffs. If we attempt to


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2198 Ind. Eng. Chem. Res., Vol. 26 , No . 11,1987overdesign the equipment beyond that needed to make theprocess operable in the face of disturbances (see Fisher e tal. (1987)), the incremental capital cost will exceed thesavings in operating costs.However, both raw materials costs and utilities costschange with time. As these costs change, the location ofthe minimum operating cost and minimum total annualcosts will change. Thus, even if the original plant corre-sponded to an optimum design (which is seldom the case),situations can arise where the expenditure of incrementalcapital can be much less than the corresponding savingsin operating costs as a particular operating variable ischanged. For these cases, we expect that the current valueof an operating variable in the existing plant will be at anequipment constraint, since this policy corresponds to themost profitable operation.We call the equipment tha t constrains changing an op-erating variable to a lower operating cost condition bot t -lenecking equipment. It should be noted that not alloperating variables are constrained, and there might besome optimization problems that involve tradeoffs onlybetween operating costs. Hence, we distinguish betweenprocess constraints (which must always be satisfied),constrained operating variables (where the constraint canbe removed by changing the existing sizes of bottleneckedequipment), and unconstrained operating variables (wherewe want to adjust th e operating variable in order to min-imize the operating costs).4. Modify Equipment Sizes for the Existing FlowSheet or a Process Alternative. In order to modifyprocess equipment sizes for a given flow sheet alternative,we proceed through the following levels.(4a) Eliminate the process heat exchangers, but retainthe heating and cooling utilities costs.(4b) Identify the most significant operating variables.(4c) Identify the piece(s) of equipment tha t constrainsthe significant operating variables.(4d)Remove the equipment constraints by adding excesscapacity until the incremental savings balance the incre-mental investment.(4e) Energy integrate the process without regard to theexisting heat exchangers.(40 Modify the new heat-exchanger network (if desira-ble) to accommodate the use of existing heat exchangers.(4g) Reoptimize the flows and heat-exchanger network(since they are coupled) if the retrofi t study is to be con-tinued.Each of these levels will be discussed below.Level 4a. Normally the heat-exchanger network de-pends more strongly on the process flows than vice versa.Moreover, if we change the process flows with a fixedheat-exchanger network, we will almost always encounterequipment constraints. Therefore, to simplify the problem,we begin by eliminating the existing heat-exchanger net-work, except for the hot and cold utility operating costs.Level 4b. In order to determine the most importantoperating variables, we use the rank-order parameter, r j ,presented by Fisher e t al. (1985). This is defined as

where r j is the rank-order parameter for variable j , x j isthe operating variable, a f i / a x j are the components of thetotal operating cost gradient, and A x j is a scaling factor.The scale factor ( k j ) s chosen on physical grounds as themaximum range of the operating variable within which weexpect to observe the optimum. Some typical values aregiven in Table 11.

Table 11. Scale Factors for ODtimization Analys isscaletvDe of oDtimization ranee of th e oDtimum factor

reflux ratio in distillationsolvent flow in gasabsorbersfractional recoveries incolumnsapproach temp in heatexchangersreactor conversionreactor tempreactor pressuremolar ratio of reactants

1.0 R / R m < 1.31.2< L / m G < 1.60.99< f < 1.03 F < AT < 25 O F0 < x < 1 or xeqobtain the range from theobtain the range from theobtain the range from the

chemistchemistchemist

0.30.40.0122 O F1 or xeq

As shown by Fisher et al. (1983, the rank-order functionseparates the optimization variables into various order-of-magnitude categories. Thus , it is a simple matter tojudge the relative importance of the optimization required;i.e., usually we can ignore the optimization variables th atare 1 or 2 orders of magnitude less important than theoperating variables with the largest rank-order values.We determine the dominant economic tradeoffs for eachoperating variable by comparing the positive and negativecontributions of the components of the gradient separately.Then, any positive terms t ha t are less than 10% of thelargest positive term are neglected in subsequent calcula-tions and the same for the negative terms. With thisapproach, we can often eliminate more than 50% of thecalculations (of course, we need to check the results oncewe have determined the new optimum conditions). Theremaining positive and negative terms correspond to thedominant tradeoffs.In order to estimate the incentive for optimizing themost important operating variables (i.e., those with thelargest rank-order parameters), we calculate the proximityparameter, Pj, defined by Fisher et al. (1985), where

Thus, the proximity parameter is confined to lie between-1 and 1. If the proximity parameter for an operatingvariable is equal to zero, that variable is at its optimumvalue; i.e., the gradient is equal to zero. However, if theabsolute value of the proximity parameter is greater thanabout 0.3, the variable is outside the region where theoptimum is relatively flat. Since the minimum operatingcost is shifted from the optimum total annual cost, asexplained above, the proximity parameters for the oper-ating variables will normally be large.Level 4c. For large values of the proximity parameter(Le., lPjl > 0.3),we change the operating variable in thedirection opposite to the sign of the proximity parameterfor that variable; i.e., if the gradient is positive, we decreasethe value of the operating variable. We continue movingin this direction until we hit an equipment constraint(which may be immediately). This identifies the piece (orpieces) of equipment which constrains the operating var-iables as we attempt to decrease the operating costs.Level 4d. Then, we calculate the incremental annual-ized capital cost required to remove the equipment con-straint and the savings in operating costs. We continueuntil the incremental costs are equal or until we hit anotherconstraint. If a second equipment constraint is encoun-


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Ind. Eng. Chem. Res., Vol. 26, NO . 11,1987 2199tered, we merely calculate the incremental, annualizedcapital costs required to remove both constrainsts andcontinue until the incremental costs balance the incre-mental savings in operating costs.Then, if time permits, we might consider wing the sameapproach for the less important operating variables. Ourprimary goals are to get some feeling for the magnitudeof possible savings from retrofitting and to see if a morerigorous retrofit study can be justified. This screening ofretrofit opportunities should provide an estimate of wherewe reach the point of diminishing returns, Le., where theeffort required to optimize an operating variable is notjustified by the potential savings.Level 4e. With the new process flows, we now developthe best heat-exchanger network.Level 4f. It is possible that the best new heat-exchangernetwork will be able t o use some of the existing heat-ex-change equipment. Hence, we modify the design to useas much existing equipment as is economical.Level 4g. In cases where potential raw material savingsare important, our initial retrofit analysis is focused onchanging the process flows. Of course, the desired amountof energy integration changes with the flows although theoptimum flows also depend on the heat-exchanger net-work. To resolve this conflict, we initially remove the heatexchangers used for energy integration, and then we re-trofit considering only the flows (this procedure willchangethe energy integration). Then, we consider the energyintegration retrofit problem. Next, we could evaluate theeffect th at the heat-exchanger network has on the newestimates of the optimum flows. However, the goal of ourscreening calculations is to see if further work can bejustified; both in terms of the simultaneous optimizationof the flows and the heat-exchanger network and morerigorous computer-aided design calculations.5. Refine the Retrofit Calculations. Our screeningcalculations provide an indication of the incentive for doingadditional work. These subsequent studies should becarried out by using computer-aided design programs.Retrofitting and Process ControllabilityWe must ensure that any changes to the structure of theflow sheet or the equipment will not make the retrofittedplant uncontrollable. The steady-state controllabilitycriterion used by Fisher and Douglas (1985) and Fisher etal. (1987) is that for every process constraint or significantoperating (optimization) variable, there exists at least onesignificant manipulative variable. The major potentialproblem with process controllability in a retrofit analysisis the elimination of manipulative variables. This isparticularly true for highly energy-integrated processes.For example, heat loads in integrated distillation columnscannot be varied independently. However, the additionof bypass valves or auxilliary heaters can be used to restoreoperability.The retrofit design should also be checked to ensure thatit can handle t he full range of process distrubances, Le.,that it is operable. In many cases, equipment overdesignto accommodate disturbances can be used to significantlydecrease the incremental operating costs required tomaintain operability when disturbances enter t he plant.These issues are discussed by Fisher et al. (1987).Example. The Hydrodealkylation of Toluene ToProduce BenzeneAs an example of the retrofit procedure, we consider thehydrodealkylation of toluene to produce benzene (HDA).This process was selected because the design of the processhas been discussed in detail in the literature. Here we

Table 111. O p t i m u m D e s ig n C o n d i t i o n s f o r t h e E x i s t i n gH D A P r o c e s sdesign variable value

reactor conversionH2 purge compositionwater cooler inlet temp, Kwater cooler outlet temp, Kbenzene recovery in produc t columnreflux ratio in product columntoluene recovery in recycle columndiphenyl recovery in recycle columnreflux ratio in recycle column

0.750.464283110.991.20.9860.8071.0

G A S RECYCLE

COMPREBDR

TOLUENEFEED

DIPHENYL4 4F i g u r e 2. HDA process flow sheet.

F i g u r e 3. Operating cost diagram for th e HDA process. Values arein $1 X 106/year. Negative values are credits.focus on the difference between the design and retrofitproblems.The original problem and an optimized design can befound in McKetta (1977). The reactions are

toluene + H2- enzene + CH42(benzene) 2 diphenyl + H2

The values of the optimized design variables are listedin Table 111, and a simple flow sheet is given in Figure 2.Since, the prices of the raw materials and energy havechanged considerably since 1967, we can examine the re-trofi t of the original design.1. Identifying the Incentive for Raw Material andEnergy Retrofitting. An operating cost diagram for theHDA process is shown in Figure 3. It is obvious from thisdiagram that incremental raw materials costs (incurred byselectivity losses to byproducts, purge losses, etc.) are muchgreater than energy costs. Thus, there is a significantincentive to look for ways of decreasing the amount oftoluene that is lost to diphenyl (and to the purge) and toreduce the amount of hydrogen tha t is lost in the purgestream. However, any changes in these flows will also


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2200 Ind. Eng. Chem. Res., Vol. 26, No. 11,1987Table IVLevel 1:Level 2:

Level 3:

Level 4:Level 4a:

Level 5:

. Design DecisionsBa tch vs. Continuous-we consider only continu ousprocesses.Input-Output Structu re of Flow Sheet

1. Should we purify the raw material streamsbefore they are fed to the reactor? If theimpurities are inert, there are no quantitativeheuristics.2. How man y product str eams will there be?Reasonable heuristics see m to be available, exceptfor the case of a reversible byproduct.3. Should a reversible byprod uct be recovered orrecycled to extinction? No quantitative heuris tic isavailable.4. D o we need a gas recycle an d a purge stre am?A quantitative heuristic seemed to be availablebefore the invention of membrane separationprocesses to s eparat e gaseous mixtures.1, How many reactor systems are required? Theheuristics seem to be reasonable. Is there anysepara tion between the reactors? Usually adecision can be m ade based on th e chemists data.2. How many recycle strea ms are there?Heuristics a re available.3. Should we use an excess of one reac tant?Normally chemists data will indicate the answer.4. I s a gas recycle compressor required? Aheuristic is available.5 . Should the reactor be operated adiabatically,with direct heating (or cooling), or is a diluent (heatcarrier) needed? Some calculations are needed touse the heuristic.6. Do we want to shift the equilibrium conversion?Calculations and judgement are required.1. Wh at is the s truct ure of the vapor an d liquidrecovery system? Heuristics are available.1.W hat is the best location of the vapor recoverysystem? No heuristics are available.2. W hat sequence of distillation columns should beused? Th e published heuristics are limited toshar p splits of ideal mixtures for a single feed, butin many cases they do not lead to the b est sequence.3. How should the light ends be removed?Calculations and judgement are required.4. Should the light ends be vented, sent to fuel, orrecycled to t he vapor recovery system?Calculations and judgement are required.5 . How should we accomplish the otherseparations? No heuristics ar e available.Heat-Exchanger Network-A design procedure isavailable (Linnhoff e t al., 1982) for developingalternative designs for heat-exchanger networks.Also, a procedure described by A ndrecovich andWesterberg (1985) can be used forenergy-integrating distillation columns.

Recycle Struct ure

Separation SystemVapor Recovery System

change the utility requirements.2. Designing a New Plant Using the Existing FlowSheet. The application of the hierarchical decision pro-cedure to the design of the HDA process has been de-scribed in detail by Douglas (1985). Th e decisions th atmust be made to develop a flow sheet are given in TableIV, and the decisions that correspond to the flow sheetshown in Figure 2 are listed in Table V. By use of currentprices, the optimum design conditions for conversion andpurge composition are significantly different than theoriginal optimum values given in Table 111. Tha t is, thenew optimum conversion is lower than the original valueso that the selectivity losses are lower. The optimum purgecomposition is also lower so that less hydrogen is lost inthe purge. These differences reflect the large change inraw materials prices over the years, (e.g., the cost of toluenein the 1967 case study was given as $0.16/gal vs. $1.30/galnow).

Table V. Process Alternatives for the HDA ProcessLevel 2 Decisions: Input-Output Structure

1.Do not purify the hydrogen feed stream.2. Recover, rather th an recycle, diphenyl so tha t the re a rethree product streams (purge, benzene product, diphenylbyproduct).3. Use a gas recycle and purge st ream .1.Use a single reactor.2. Use a gas (H 2and CH,) and a liquified (toluene) recyclestream.3. Use a 5 / 1 H2/a rom atics ratio to prevent coking-assumingthis to be a design constraint (although it could be formulatedas an optimization problem).4. A gas recycle compressor is needed.5 . Operate the reactor adiabatically.6. Dont consider equilibrium effects.1. All separations should be by distillation.2. Use the direct sequence of simple columns-probablycomplex columns should be used.3. Remove light end s in stabilizer.4. Th e light ends are sent to fuel-no vapor recovery systems.Level 5 Decisions: Energy Integration-there are nume rousalternatives.3. Quick Screening of Process Alternatives. Th ehierarchical decision procedure of Douglas (1985) makes

it a simple matter to generate a list of process alternatives,Le., we merely change the decisions given in Table IV. Aprocedure for quickly screening these alternatives for thedesign of a new HDA process has been presented byDouglas and Woodcock (1985). The results indicate thatthere is a large incentive for decreasing the selectivity ossesby recycling the diphenyl to extinction and that there isa large incentive for recovery of hydrogen from the purgestream.Of course, the design problem is different from th e re-trofit problem, because in the design problem we alwayscalculate the equipment sizes as a function of the designvariables, whereas in the retrofit problem we fix the valuesof the equipment sizes. If diphenyl is recycled to extinction(i.e., to its equilibrium value), then there are no selectivitylosses, the optimum conversion of toluene changes fromx = 0.675 to x = 0.997, and the total annual cost of theprocess is reduced from $4.73 X 106/year to $3.57 X106/year (Terrill and Douglas, 1987). In a retrofit analysis,we must determine whether the existing equipment cantolerate a large change in flows corresponding to this newvalue of conversion, and if not we must consider the ad-dition of new equipment.4. Modifying Equipment Sizes for the ExistingFlow Sheet or a Process Alternative. Obviously, it isnecessary to evaluate each of the process alternatives froma retrofit viewpoint. However, for the purposes of illus-trating our procedure, we only consider changes to theexisting flow sheet. Normally, we would evaluate this typeof change first, because we expect that the fewest numberof changes would be required (Le., the lowest incrementalinvestment).The first step in examining equipment modifications isto remove the feed-effluent heat exchanger in Figure 1from the process (i.e., wewillperform an energy integrationanalysis after we have modified the flows). We retain thefurnace and partial condenser, however, so that we canconsider their operating costs (although we do not considerequipment constraints in these units a t this stage of t heanalysis). We retain these operating costs because someenergy from hot and cold utilities will always be required.Next we determine the significant operating variablesand the economic tradeoffs for these optimizations bycalculating the rank-order functions and the proximity

Level 3 Decisions: Recycle Stru ctur e

Level 4b Decisions: Liquid Separ ation Systems


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Ind. Eng. Chem. Res., Vol. 26 , No . 11, 1987 2201Table VI. Gradients of Cost Functions at Base-CaseConditions

f, afiiax a m y p H afiiaFm afiiaRppfurnace fuelcompressor powerproduct column utilitiesbenzene losses in prod .purge lossesselectivity lossesrecycle column utilitiescoling water (flashX j

col.

d r u m )Ax j

-1134-12500

-3115577-3820

-1156-20100

7135000

0 00 00 1540 -154-0.5 00 00 0

0 00.75 0.46 150 1.270.1 0.1 50 0.1750 850 50 300.48 0.77 0 0

parameters. These values are shown in Table VI.The conversion per pass of toluene and the hydrogenpurge composition are clearly the most important opti-mization variables. Their rank-order parameters are atleast an order of magnitude larger than the remainingvariables. Also, thei r large proximity parameter valuessuggest that smaller values for both conversion and purgecomposition should lead to significant reductions in op-erating costs, but these changes might not be attainablewith existing equipment capacities.Curiously, it can be shown that the same piece ofequipment constrains the optimum value of both of thesevariables. Operation at lower hydrogen purge compositionswould require higher gas recycle flows to satisfy the 5 : lhydrogen-to-toluene ratio requirement at the reactor inlet.Thus , the gas recycle compressor capacity constrains theoptimum purge composition.Operation at lower reactor conversions would naturallyrequire higher toluene recycle flows. An unusually highdesign value for the reflux ratio in the recycle column forthe original design ensures that sufficient vapor capacityis available to handle the increased toluene recycle load.However, in this case a larger gas recycle flow would alsobe required to maintain H2 /T = 5. Thus, the optimumreactor conversion is also constrained by the gas recyclecompressor capacity.Because the compressor capacity constrains the twomost important optimization variables, it provides ourmost likely candidate for retrofitting. Both variablesprimarily trade off raw material costs vs. heating costs forincreased recycle flows. Thus, energy integration aimedat reducing these costs may also be appropriate.The cooling water flow rate and the reflux ratio in theproduct column are less important optimization variables.Moreover, these variables are able to attain their optimumvalues with the existing equipment (Pj= 0 for each).The cooling water flow rate primarily trades off utilitiescosts vs. purge losses of benzene and toluene. The onlyretrofit policy possible to reduce these costs would be toincrease the area of the partial condenser preceding theflash drum. The reflux ratio in the product column pri-marily trades off product losses to the bottoms stream vs.heating and cooling costs for an increased reflux ratio.Possible retrofit policies would be energy integration toreduce the utilit ies requirements or increased number oftrays in the stripping section to improve the product re-covery. Implementing this last policy is unlikely for anexisting column.Now that the most attractive retrofit policies have beenidentified, the cost of equipment modifications must beincorporated into the optimization analysis. In particular,we must specify whether the capacity of a piece of

- 1

43 tL56I I 1 I I0 10 20 30 40 50

COMPRESSOR OVERDESIGN (%IFigure 4. Estimation of the optimum retrofit policy for the gasrecycle compressor.equipment can be increased incrementally (via auxiliaryexchangers, etc.) or whether a new unit must be introducedto replace an existing one. Only by comparing the po-tential decrease in opeating costs with the incrementalcapital costs incurred can we decide if a retrofit policy isbeneficial.Gas Recycle Compressor. In order to increase the gasrecycle flow rate, we assume that we can introduce anothercompressor in parallel with the existing unit. As the ca-pacity of this new compressor increases, the constrained(optimum) values for both purge composition and reactorconversion decrease. This tradeoff between the addedannualized compressor capital cost and the total annualoperating cost for the process is shown in Figure 4.Based on a capital charge factor of 113 per year (ofcourse, different companies use different capital chargefactors), the approximately optimum retrofit policy is toinstall an additional gas recycle compressor with 56% ofthe capacity of the existing unit. The incremental capitalcost for the compressor just balances the reduction inoperating costs (where the slopes in Figure 4 are equal andopposite in sign). Thi s policy incurs a capital cost of$55000/year but reduces the total operating costs by about$480000/year. Of course, the greatly increased gas andliquid recycle flows would also saturate the furnanceheating capacity. Again, this conflict will be addressedafter a revised energy integration analysis has been com-pleted.

Flash Drum Cooler. In order to drive the flash drumtemperature closer to the cooling water tempera ture , wecould introduce another exchanger in series with the ex-isting partial condenser. The optimum exchanger areawould primarily trade off the associated capital cost vs.the purge losses of benzene and toluene. That is, the flashdrum temperature has only a small effect on the remainingoperating costs. As shown in Figure 5, the approximatelyoptimum retrofit policy is to install an additional ex-changer with 1.5times the area of the existing unit. Thispolicy incurs a capital cost of $75000fyear but reduces thetotal operating costs by about $410 000fyear.Energy Integration Analysis. The energy integrationfor the original flow sheet, Figure 2, consists solely of afeed-effluent heat exchanger (FEHE) around the high-


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2202 Ind. Eng. Chem. Res., Vol. 26, No . 11,1987

3

,oooi

I , , I ,

/ /

I 1 I I 1 I0 100 200 300

CONDENSER OVERDESIGN (%)F i g u r e 5. Estimation of the optimum retrofit policy for the partialcondenser preceding the flash drum.

12001

/-01 I I 1 I 10 20 40 60 80 1 0 0 120

HEAT DUTY (lo6 Btu/hr)F i g u r e 6. Temperature-enthalpy diagram for the HDA process atbase-case operating conditions.T a b l e VII. C o m p a r i s o n of t h e B a s e -C a s e a n d R e t r o f i tO p e r a t i n g C o n d i ti o n s

original retrofittedreactor conversion 0.75 0.69purge composition 0.46 0.35recycle col. reflux ratio 1.27 1.36cooling water flow ( % of design) 150 12 0furnace duty ($ X lo6B t u / h ) 11.9 11.9

temperature reactor. Figure 6 shows the temperature-enthalpy diagram for the original operating conditions andequipment sizes. Th e incentive for additional energy in-tegration is shown quantitatively by the separation of theheating-and cooling curves. Th e new optimum operatingconditions based on the above retrofit analysis are givenin Table VII. The resulting temperature-enthalpy dia-gram is given in Figure 7.On the basis of this information, we can make two im-portant observations: (1) the incentive for energy inte-gration is much greater for the new process flows and (2 )

r I \ I

I -b 3

68I 1 1 1 I0 20 40 60 80 100 120

EXCHANGER OVERDESIGN (%)F i g u r e 8. Estimation of the optimum retrofit policy for the feed-effluent heat exchanger.the required furnace load is 60% greater than the designvalue. Thus, one goal of any additional energy integrationshould be to reduce the required furnace load and to re-store operability. This can most easily be achieved byincreasing the area of the existing FEHE. Other integra-tion strategies (suchas hermal coupling of the distillationcolumns) should also be investigated.In a detailed study of heat-exchanger networks for thedesign of an HDA process, Terrill and Douglas (1987)found tha t a feed-effluent exchanger alone provides closeto the optimum design. Thus, we might begin our retrofitof energy integration by estimating the optimum amountof overdesign for the FEHE. Th e tradeoffs involved areshown in Figure 8. The approximately optimum retrofitpolicy is to install an additional exchanger with 68% ofthe area of the existing one. This policy incurs a capitalcost of about $160 000/year, while reducing the total op-erating costs by about $250 000/year.This optimum retrofit policy for the FEHE reducesthe required furnace load to within 10% of its design value.If this is still not tolerable, additional overdesign of the


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Ind. Eng. Chem. Res., Vol. 26, No. 11,1987 2203An interactive program, PIP (Process Invention Proce-dure), tha t contains both heuristics and algorithms thatwill accomplish the first three steps in the procedure, aswell as step 4b, has been described by Kirkwood et al.(1987). This program has a different achitecture than aconventionalAI expert system because the knowledge basehas been arranged as a hierarchical plan (Sacerdoti, 19741,the knowledge has been precedence ordered a t each levelin the hierarchy, both a qualitative (heuristics) and aquantitative (algorithms) knowledge base have been in-

cluded a t each level in the hierarchy, and th e heuristicsthat are used to fix the structure of the flow sheet and toidentify the process alternatives are also used to select theappropriate subroutines for the design calculations.A code for screening retrofit opportunities (for the same,limited class of processes considered by PIP) can be de-veloped by adding a set of shortcut performance modelsto PIP and by adding the logic required to recognizeequipment constraints and to switch from performancemodels to design models when constraints are encountered,to plot the incremental and annualized capital costs andoperating cost savings as the operating variables arechanged, and to retrofit heat exchanger networks after theflows have been changed. An attempt to devlop a code forthis purpose is currently under way.ConclusionsIt is possible to identify three types of retrofitt ingproblems: (1)process "debottlenecking" to increase pro-duction capacity, (2 ) expansion of equipment capacity toease constrained operation, and (3) modification of un-constrained equipment to reduce operating costs. Heat-exchanger network retrofit procedures concern themselvesprimarily with this last item but do not normally addressthe remaining issues.The systematic retrofit procedure developed in thispaper considers each of these problems in the order given.Bottlenecking equipment can be identified by using apublished flow sheet decomposition scheme. Quantitativeparameters used t o characterize the design optimizationproblem for new processes can be used to evaluate theincentive for additional retrofitting. The procedure canbe used as the basis for an interactive code for screeningretrofit opportunities for petrochemical plants.The procedure has been applied to several publishedcase studies, where the goal of retrof itting was to reduceoperating costs at current production rates and operatingconditions (a common case). In each of the case studiesthere were large selectivity losses of raw materials to by-products. The results indicate tha t there is a significantincentive to modify the process operating conditions if rawmaterial costs have changed significantly from the originaldesign conditions. The new stream flow rates and heat-ing/cooling loads also increase the incentive for energyintegration. Nonetheless, the final decrease in annualizedcosts is an order of magnitude larger than is possible withenergy integration alone.Acknowledgment

We are grateful to the National Science Foundation forpartial support of this work under Grant CPE-8105500.Nomenclaturef = fractional recovery (Table I)f i = cost functionF,, = flow rate of cooling water to flash drum coolerFEHE = feed-effluent heat exchangerHDA = hydrodealkylation of toluene to benzene processP, = proximity parameter for optimization variable x j

FE HE could be used to reduce the furnance load further.As illustrated in Figure 8, this would result in only a smallincrease in the total annualized costs for the process. Wecan use calculations of this type in order to make a pre-liminary assessment of the incentive for undertaking amore detailed energy integration retrofit study. Of course,if our screening calculations indicate tha t additional effortcan be justified, we want to repeat our analysis by usingrigorous calculation procedures and a more detailed energyintegration retrofit procedure; see Tjoe and Linnhoff(1986).A much more detailed energy integration analysis forth e design of this process has been conducted by Terrilland Douglas (1987). They found that a feed-effluent heatexchange alone provides a nearly optimum heat-exchangernetwork for a new process design. Tha t is, more complexnetworks provide only minimal (6%)eductions in thetotal annualized costs for this process. For the retrofitanalysis, the flow sheet is much less flexible with morecomplex networks, so we assume that a feed-effluent ex-changer is close to optimum.Summary of the Initial Retrofit Analysis. Theproposed retrofit policies for the HDA process incurreda totalof $290000/year in capital costs, while reducing thetotal operating costs by $1 140 000/year. The policiesprovide a 130% return on a capital investment of $870000(o r a 9-month payback period). If this is sufficient jus-tification to proceed further with the retrofit study, thenwe must iterate between the flow optimizations and theheat-exchanger network optimization. This provides agood starting point for computer-aided design calculations.We should note that energy integration alone reducedthe total annualized costs for the process by about$90000/year, accounting for only 10% of t he savings forthe complete retrofit analysis. While this result is specificto this example, it nonetheless points out the importanceof considering both flow optimizations and heat-exchangernetwork alteratives during the retrofit analysis. Clearly,the "best" heat-exchanger network for the original con-ditions, Figure 6, would differ significantly from anynetwork based on Figure 7. As will often be the case, theproposed retrofit policies for the HDA plant increase thetotal utility heating and cooling, as well as increase theamount of process/process heat exchange.Other AlternativesOf course the retrofi t policy described above may notbe the best one to choose. Thus, we also need to considerthe other process alternatives that we identified earlier.However, the general approach of identifying the signifi-cant operating variables and modifying equipment sizesis the same.Developing an Interactive Computer Code forScreening Retrofit OpportunitiesThe retrofit procedure that we present in Table I iscomputationally intensive; i.e., a large number of calcu-lations are required, and the number of production rulesas compared to the number of algorithms is much less thanin the synthesis procedure described by Douglas (1985).Hence, the proceudre is not well suited for current AI(Artificial Intelligence) expert system shells which handleonly qualitative knowledge (e.g., MYCIN, KEE, etc.).Similarly, the calculations are tedious if conventionalsimulators such as PROCESS, ASPEN, DESIGN 2000,Chem CAD, etc., are used, because it is necessary to switchfrom a performance to a design model (and to recompilethe program) every time that an equipment constraint isencountered as the operating variables are changed.


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2204 Ind. Eng. C h e m . Res. 1987,26, 2204-2211r . = rank-order parameter for optimization variable x ,R = reflux ratioR , = minimum reflux ratioRPc = reflux ratio for the product columnx = reactor conversionxx;= optimization variableYpH = hydrogen purge compositionLiteratur e CitedAndrecovich, M. J.; W esterberg, A. W. A Simple Synth esis Method

Based on Utility Bounding for Heat-Integrated DistillationSequences, AIC hE J . 1985, 31 , 363.B o la n d , D . ; H in d ma rs h , E . H e a t E x c h a n g e r N e tw o rkImprovement, CE P 1984, 8 0 ( 7 ) ,47.Douglas, J. M. A Hierarchical Decision Procedure for ProcessSynthesis, AIChE J . 1985, 31, 353.Douglas, J. M.; Woodcock, D. C. Cost Diagrams and the QuickScreening of Alternatives, Ind. Eng. Chem. Process Des . Deu.1985, 24 , 970.Fisher, W. R.; Doherty, M. F.; Douglas,J. M. Evaluating SignificantEconomic Tradeoffs for Process Design and Stead y-Sta te ControlOptimization Problems, AIChE J . 1985, 31, 1538.Fisher, W. R.; Douglas, J. M. E valuating Process Operability a t thePreliminary Design Stage, Comp. Chem. Eng . 1985, 9, 499.Fisher, W. R.; Doherty, M . F.; Douglas, J.M. The Interface BetweenDesign and Control, Ind. Eng. Chem. Res. 1987, in press.Kirkwood, R. L. ; Locke, M. H.; Douglas, J. M. An Expert Systemfor Synthesizing Flowsheets and Optimum Designs, Comp.Chem. Eng. 1987, in press.

= equilibrium reactor conversion (Table I)

Linnhoff, B.; Vredeveld, D. R. Pinc h Technology Has com e of Age,C E P 1984, 80(7), 33.Linnhoff, B.; Townsend, D. W.;Boland, D.; Hewitt, G. F.; Thom as,B. E. A.; Guy, A. R.; Marsland, R. H.; A User Guide on ProcessIntegration for the Efficient Use of Energy, Ins t . Chem. Eng.1982, 1.McKetta , J. J., Ed. Encyclopedia of Chemical Processing andDesign, Marcel Dekker: New York, 19 77; Vol. 4, p 182.Sacerdoti, E. D., Plan ning in a Hierarchy of Abstraction Spaces ,Artif. Intelligence 1974, 5, 115.Steinmetz, F. J.; Chaney, M. D. Tota l Pla nt Energy Integratio n,Presented at the Spring National AIChE Meeting, Houston,Mar ch 24-28, 1985; Pap er 88d.Terrill , D. T.; Douglas, J. M. H eat Exchanger Network Analysis. 1.Optimization, Ind. Eng. Chem. Res. 1987, 26, 685.Tjoe, T. N.; Linnhoff, B. Using Pinch Technology for ProcessRetrofit, Chem. Eng. 1986, April 28, 47.Townsend, D. W.; Linnhoff, B. Surface Area Targets for Hea t Ex-changer Networks, Annual Meeting of the Institute of ChemicalEngineers , Bath , U.K., April 1984.Westerberg, A. W. The Role of Expert System Technology inDesign, Paper presented at the International Symposium onChemical Reaction Engineering, Philadelphia, May 18-21, 1986;ISCRE 9 .Witherell, W. D.; Linnhoff, B. Pinc h Technology Retrofit: A Com-plex Industrial Application, Presented at the Spring NationalAIChE Meeting, Housto n, March 24-28, 1985; Paper 88b.

Received for review August 15, 1986R e v is e d ma n u s c r ip t r e c e ive d J u n e 3, 1987Accepted Ju ne 27 , 1987

Cocracking and Separate Cracking of Ethane andNaphthaPatrick M. Plehiers and Gilbert F. Ftoment*Labora tor ium voor Pe trochemische Techniek , Ri jksun ivers i te i t te Gent , B-9000 Gent , Be lg iumThis pa per presents experimental data on the thermal cracking of a naphtha-ethane mixture in apilot plant, under conditions representative of industrial operation. Th e effects of th e interactionbetween ethane and naph tha on the ethane conversion and kinetics a nd on th e product distributionare investigated. When naphtha, ethane, and th e mixture are cracked under an equal molar dilution,the cocracking yields and selectivities can be quantita tively predicted from the separate crackingdata, except for hydrogen, methane, an d high molecular weight products. Th e combined effectsof cocracking an d partial pressure, occurring when naphth a, ethane, an d the mixture are crackedunder an equal weight dilution, are such th at, if maximum olefins selectivities are desired, separatecracking is to be preferred. Some aspects of coke formation are addressed as well.

IntroductionLiterature data on naphtha-ethane pyrolysis are scarceand very incomplete. Most authors base their conclusionsupon one single data point. de Blieck and Goossens(1971a,b) observed that cocracking with naphtha increased,for identical reaction conditions, the ethane conversion.From a comparison of typical product distributions forcocracking and separate cracking, Mol (1981) concludedthat the interaction between ethane and naphtha in thecracking of a mixture containing about 24% by weightethane leads to an enhanced ethylene selectivity, enablinga 2YO savings in naphtha consumption. Propylene andbutadiene selectivities, however, were markedly reduced.Nowowiejski et al. (1982) cracked naphtha with ethane inan industrial millisecond furnace. In addition to themillisecond effects, they found the methane and ethyleneyields t o be favored by cocracking; the C,+ and butenesyields were found to be reduced. No accelerat ing effectof the naphtha on the ethane cracking was noticed in thiscase. A considerable influence of the naphtha compositionon the deviations of the cocracking yields from additivity

0888-5885/87/2626-2204$01.50/0

was observed, but no further details are given on this issue,however.Clearly, until now, no thorough study of naphtha-ethanecocracking has been published. The work reported in thepresent paper aimed at a better understanding of the in-teraction between naphtha and ethane during pyrolysis.I t was investigated how the interaction alters the overallethane cracking kinetics and the product distributions.Experiments were performed under conditions close tothose encountered in industrial practice.Statement of the Problem

The aim of a study on cocracking of hydrocarbons is tocompare the cocracking yields with those resulting fromthe mixing of the effluents of separate cracking. Quiteoften in the literature, identical reaction conditions arechosen as a basis for this comparison. Since differenthydrocarbons demand completely different operatingconditions, comparing yields at equal conditions does notseem very appropriate. Moreover, identical reaction con-ditions do not guarantee the naphtha and ethane con-0 1 9 8 7 Society

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Screenging of Process Retrofit Alternatives

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