Reactor Network Synthesis Vinyl Chloride Proces

Computers and Chemical Engineering 23 (1999) 479–495

A case study for reactor network synthesis: the vinyl chlorideprocess

A. Lakshmanan, W.C. Rooney, L.T. Biegler *,1

Chemical Engineering Department, Chemical Mellon Uni6ersity, Pittsburgh, PA 15213, USA

Received 11 February 1998; received in revised form 3 July 1998

Abstract

A key objective of the integrated reactor network synthesis approach is the development of waste minimizing process flowsheets(Lakshmanan & Biegler, 1995). With increasing environmental concerns in process design, there is a particularly strong need tomaximize conversion to product and avoid generation of wasteful byproducts within the reactor network. This also avoidsexpensive treatment and separation costs downstream in the process. In this study, we present an application of the mixed integernonlinear programming (MINLP)-based reactor network synthesis strategy developed by Lakshmanan and Biegler (1996a). Herewe focus on applying these reactor network synthesis concepts to the vinyl chloride monomer production process. Vinyl chlorideis currently produced by a balanced production process from ethylene, chlorine and oxygen with three separate reaction sections:oxychlorination of ethylene; direct chlorination of ethylene; and pyrolysis of ethylene dichloride. The hydrogen chloride producedin the pyrolysis reactor is used completely in the oxychlorination reactor. Byproducts such as chlorinated hydrocarbons andcarbon oxides are generated by these reaction sections. These are studied using reaction kinetic models for the three reactionsections. The case study results in optimal reactor networks that improve the conversion of ethylene to vinyl chloride andminimize the formation of byproducts. These results are used to generate an improved flowsheet for the production of vinylchloride monomer. Moreover, an overall profit maximization, that includes the effect of heat integration, is presented and a setof recommendations that improve the selectivity of vinyl chloride production are outlined. Finally, the optimal reactor structures,overall conversion and annual profit are shown to be only mildly sensitive with respect to small changes in the kinetic parameters.© 1999 Elsevier Science Ltd. All rights reserved.

1. Introduction

In most chemical processes the selective conversion ofraw materials in the reactor determines the compositionand the amount of waste products produced, and thereactor network has a significant influence on the recyclestructure and downstream processing steps. Hence, re-action systems and reactor design often determine thecharacter of the flowsheet and better performance of thereactor system improves the performance of the entireprocess. In the past decade, several researchers haveworked on problems of reactor network synthesis andreactor design. This problem involves nonlinear reactionmodels, uncertain rate laws and numerous possiblereactor types and networks. Hence, researchers in thisarea have achieved limited success. Motivated by thesethoughts Lakshmanan and Biegler (1996a) proposed a

reactor network synthesis strategy that combines princi-ples developed for MINLP optimization of prepostu-lated reactor superstructures and geometric techniquesfor establishing the reactor network.

In particular, geometric techniques construct an at-tainable region, which represents the limits that can beachieved by the processes of reaction and mixing. Thisregion includes the performance of any reactor networkstructures in terms of conversion of raw materials todesired products and the amount of waste products thatare produced. Glasser, Crowe and Hildebrandt (1987)and Hildebrandt, Glasser and Crowe (1990) proposedconstructive techniques to map the attainable region forreactor network synthesis problems in two or threedimensions. These techniques are useful to develop newreactor structures and to understand the reactor net-work synthesis problem, but it is difficult to map theseregions in more than three dimensions by using thegeometric approach. On the other hand, the optimiza-tion based approach of Lakshmanan and Biegler

* Corresponding author.1 Present address: Aspen Technology, Cambridge, MA02141, USA.

0098-1354/99/$ - see front matter © 1999 Elsevier Science Ltd. All rights reserved.

PII: S 0 0 9 8 - 1 3 5 4 ( 9 8 ) 0 0 2 8 7 - 7

A. Lakshmanan et al. / Computers and Chemical Engineering 23 (1999) 479–495480

(1996b) overcomes this dimensionality problem andallows for easy integration of the reactor network syn-thesis problem with other process subsystems such asseparation, heat integration and waste treatment. Thispaper applies this integrated reactor network synthesisstrategy to the vinyl chloride monomer productionprocess.

Vinyl chloride monomer is one of the largest com-modity chemicals produced in the United States. One ofthe first patented processes to produce vinyl chlorideinvolved reacting hydrogen chloride and acetylene inthe presence of a mercuric chloride catalyst. Nowadays,it is produced commercially by pyrolytic decompositionof ethylene dichloride (EDC). EDC may be producedby the direct chlorination of ethylene or oxychlorina-tion of ethylene in the presence of oxygen and hydrogenchloride. Pyrolysis of EDC produces vinyl chloride andan equal amount of hydrogen chloride as a co-product,which is one of the reactants in the oxychlorination ofethylene. Hence, the hydrogen chloride produced in thepyrolytic reactor is recycled to the oxychlorination re-actor where it is completely consumed. The componentprocesses of direct chlorination, EDC pyrolysis andoxychlorination are combined to develop a balancedprocess for the production of vinyl chloride fromethylene and chlorine with no net consumption orproduction of hydrogen chloride. Apart from the mainproducts, a variety of byproducts are produced in thesereactors. Many of these byproducts are now classifiedas hazardous and are expensive to treat and dispose.Hence, there is a significant need for performance im-provement in the reactor sections. This paper focuseson developing optimal reactor configurations for theprocesses of direct chlorination and oxychlorination ofethylene and pyrolysis of EDC.

In the next section we present a brief literature surveyof the area of reactor network synthesis. Section 3outlines the mixed integer nonlinear programming(MINLP) based reactor network synthesis algorithm(Lakshmanan & Biegler, 1996a). Section 4 presents adetailed discussion of the vinyl chloride process andSection 5 applies the algorithms outlined in Section 3 tothe vinyl chloride monomer production process and theresults of this case study are presented. A candidateflowsheet and a set of recommendations that improvethe conversion of ethylene to vinyl chloride product arealso proposed. Section 5 also presents an economicoptimization of the process flowsheet and highlightschanges in the reactor structure when the objective is tomaximize the profit rather than conversion to the de-sired product. Here a brief network sensitivity analysisto kinetic parameters is also included. In addition, theprocess flowsheet is energy integrated and this results insignificant cost savings. Finally, Section 6 summarizesthe results and presents some ideas to further improvethe candidate flowsheet presented here.

2. Literature review on reactor network synthesis

Academic research to develop systematic optimiza-tion techniques to synthesize reactor networks has itsroots in the work of Aris (1960, 1961), Horn (1964),Horn and Tsai (1987) and Jackson (1968). Recentapproaches to reactor network synthesis may beclassified under two broad categories: superstructureoptimization and attainable region targeting. In super-structure optimization, a fixed network of reactors isfirst postulated based on heuristics and physical in-sights. The optimal subnetwork which maximizes theperformance index is derived from this superstructureusing mixed integer nonlinear programming (MINLP)techniques (see Viswanathan & Grossmann, 1990).In attainable region targeting, an attempt is made tofind an achievable bound on the performance index ofthe system irrespective of the actual reactor configu-ration. A general functional representation is usedto model the entire variety of reaction and mix-ing states. Bounds are then derived based on limitsposed by reaction kinetics on the space of concen-trations achievable by the processes of reaction andmixing.

2.1. Superstructure optimization

Researchers have attempted to develop optimal reac-tor networks by postulating a variety of reactor super-structures. Chitra and Govind (1981, 1985) studiedPFR systems with a recycle stream from an intermedi-ate point along the reactor and optimized the recycleratio and the point of recycle. This reactor system couldmodel PFRs, CSTRs and recycle reactors. Later, theyclassified different reaction mechanisms and postulateda superstructure consisting of a series of recycle reactorswith bypass streams and heat exchangers at the reactorinlets. Kokossis and Floudas (1990, 1991) consid-ered a large superstructure of isothermal and non-isothermal networks of PFRs and CSTRs, by modelinga PFR as the limit of a large number of sub-CSTRs(CSTRs of equal size). Their MINLP formulation wasfree of differential equations. Later, the technique wasextended to handle stability of reactor networks andintegration with recycle systems. Recently, Smith andPantelides (1995) proposed a synthesis technique forreaction and separation networks, using detailed unitoperation models. Complete connectivity among theunits, both forward and recycle, was assumed in thesuperstructure. However, these approaches may be sub-optimal since the solution obtained is only as rich asthe initial superstructure chosen and it is difficult toensure that all possible networks are included in theinitial superstructure.

A. Lakshmanan et al. / Computers and Chemical Engineering 23 (1999) 479–495 481

2.2. Attainable region targeting

Targeting for reactor network synthesis is based onthe concept of the attainable region (AR) in concentra-tion space, first suggested by Horn (1964). The attain-able region is the convex hull of all of theconcentrations that can be achieved starting from thefeed point using only reaction and mixing. Recently,Glasser et al. (1987), Hildebrandt et al. (1990) devel-oped a geometric technique to map the entire region inthe concentration space that is attainable from a givenfeed concentration by the processes of reaction andmixing. Alternate plug flow reactor (PFR) and continu-ous stirred tank reactor (CSTR) trajectories were drawnto cover the attainable region and derive an optimalreactor network. However, these geometric techniquesare difficult to apply beyond three dimensions. Toovercome this limitation, Omtveit, Tanskanen and Lien(1994a) considered higher dimensional problems by us-ing the principle of reaction invariants (Fjeld, Asb-jornsen & Astrom, 1974) and by introducing systemspecific constraints (for example, coking) to project theproblem into two dimensions. On the other hand, Bal-akrishna and Biegler (1992a,b) adapted the geometrictechnique for targeting to a mathematical programmingbased framework. The Balakrishna and Biegler al-gorithm is a constructive technique which optimizes theobjective at each stage. However, this is a sequentialapproach which may exclude some structures consid-ered by the AR approach.

In addition to developing techniques purely for reac-tor network synthesis, several researchers have devel-oped systematic approaches to simultaneouslysynthesize the entire process. Pibouleau, Floquet andDomenech (1988) considered a network of CSTRs anddistillation columns, with limited connectivity betweenthe reaction section and the separation section. Kokos-sis and Floudas (1991) considered synthesis of isother-mal reactorseparator-recycle systems with almostcomplete connectivity between the reactor network(consisting of CSTRs) and the separation network.Omtveit, Wahl and Lien (1994b) and Glasser et al.(1987) also applied the geometric targeting technique tothe synthesis of reactor-separator systems. Balakrishnaand Biegler (1993) and Lakshmanan and Biegler (1995,1996a) integrated the synthesis of reactor networks withsimultaneous flowsheet synthesis. They demonstratedtheir technique on a simplified process flowsheets forthe production of allyl chloride and the Williams Ottoprocess flowsheet. Lakshmanan and Biegler (1996b)extended this approach to consider simultaneous reac-tion, mixing and separation by integrating the reactornetwork synthesis algorithm with mass exchange net-work synthesis. Along the same lines, Papalexandri andPistikopoulos (1996) proposed a generalized modelingframework based on mass/heat transfer principles and

the representation potential of this framework wasdemonstrated with examples on specific unit operationslike reactive distillation.

3. Algorithm for reactor network synthesis

Lakshmanan and Biegler (1996a) proposed a combi-nation of superstructure and attainable region targetingtechniques to synthesize optimal reactor networks. Thisapproach incorporates attainable region propertiesderived from Feinberg and Hildebrandt (1997). Someof these properties are: (i) the boundary of the attain-able region for reaction and mixing consists of plugflow reactor trajectories and straight lines; (ii) recyclestreams need not be considered to map the boundary;and (iii) only continuous stirred tank reactor (CSTR),plug flow reactor (PFR) and differential sidestreamreactor (DSR) trajectories make up the entire boundaryof the attainable region.

The reactor network synthesis technique considersmultiple reactor paths at each stage by targeting theattainable region using reactor modules. A reactormodule consists of a differential sidestream reactor(DSR) and a CSTR. The sidestream entering the DSRin the Ith reactor module could be the feed stream tothe DSR or it could be from any of the reactor moduleexits or any combination thereof (i.e. from any pointwithin the candidate AR). The DSR sidestreamflowrate is determined in the optimization and may beset to zero, hence, a separate PFR model need not beconsidered in the reactor module. A typical reactormodule is shown in Fig. 1. A binary variable is associ-ated with each reactor path in a module. If the DSR ischosen in the Ith reactor module, the binary variableassociated with it, YId, is set equal to 1.

Additional reactor modules are added successivelyand the superstructure is developed in a constructivemanner by solving a sequence of improving MINLP’s.The attainable region properties ensure that theMINLP is compact. The steps involved in the construc-tive stagewise algorithm are: (i) Solve a segregated flowmodel to obtain a lower bound on the solution. For

Fig. 1. Ith Reactor module.


yield or selectivity objectives and fixed feeds and tem-peratures, this formulation leads to a linear programwith a global solution; (ii) Initialize the first reactormodule with the solution obtained from step 1 andoptimize it with respect to a specific objective. The inletconditions to this reactor module are the feed condi-tions. This yields an initial target to the attainableregion. In addition, this solution does not eliminateother network paths; the binary variables associatedwith each path are merely set to one or zero; (iii)Extend the reactor module with an additional reactormodule. The feed to the second module is the exit ofthe first module or a combination of the fresh feed andthe exit of the first module. If the extension improvesthe objective then further extensions need to be consid-ered, else the optimal network is assumed to have beenfound; (iv) Ensure that the feed to the Ith reactormodule extension may be the exit of any one or acombination of the exits of the previous I−1 modulesor the initial feed; and (v) Account for bypass streamsby ensuring that the exit from the Ith reactor moduleplus the bypasses from the exits of any one or acombination of the previous (I−1) reactor modules,forms the inlet to the (I+1) th reactor module.

The reactor network synthesis problem is formulatedas an MINLP as in (Eq. (P1)):

maxJ(Xexit, t)q, f, T

(P1)

XCSTR exitI=R(XCSTRI, TCSTRI)tcI+XIf (1)

dXDSRI/daI=R(X(aI), T(aI))

+ (q(aI)QsideI/Q(aI))(XsideI−X(aI)) (2)

XDSROI=XIf (3)

XDSR exitI=&�

0

f(aI)X(aI) daI (4)

1=&�

0

f(aI) daI (5)

1=&�

0

q(aI) daI (6)

t=&�

0

& a

0

(q(a I%)QsideI/QexitI− f(a I%)) da I% daI (7)

TDSR exitI=&�

0

f(aI)T(aI) daI (8)

FIf= %I−1

k=0

FkI−1 (9)

FIf XIf= %I−1

k=0

FkI−1Xk (10)

XIf=XIc in=XId in (11)

FIf=FIc+FId (12)

FIf XI exit=FIcXCSTR exitI+FIdXDSR exitI (13)

1=YIc+YId (14)

FIf XI= %Nmod

j=I

FIj XI (15)

Xexit=XNmod exit 05FIc5UYIc 05FId5UYId

YIc�{0, 1} YId�{0, 1}

Here, all flowrates are volumetric and concentrationsare based on mass, with

is the flowrate at the inlet of the IthFif

reactor module.are the flowrate and concentrationFkI−1, XkI−1

from the exit of the kth module whichis an inlet stream to the Ith module(k=0, I−1)

Xif is the concentration at the inlet to theIth reactor module.are the flowrates of the streams passingFIc and Fid

through the CSTR and DSR in the Ithreactor module. The product of FIf*XIf

is in mass flow units.are the concentration at the inlet andXIc in, Xic

exit of the CSTR and DSR XId in, Xid,respectively in the Ith reactor module.is the concentration at the exit of theXI exit

Ith reactor moduleYIc, YId are the binary variables associated with

the CSTR and DSR in the Ith reactormodule.

XI side sidestream composition of the Ith DSRreactor module. It may be specified ordetermined from any stream in the re-actor network.sidestream flowrate of the Ith DSR re-QI side

actor module.

Eqs. (1)–(8) constitute the reactor module at stage Imade up of a CSTR (Eq. (1)) and a DSR (Eqs.(2)–(8)). The inlet conditions to the module and theindividual reactors are given by (Eqs. (9)–(12)). Eq.(13) represents the exit from the Ith module and Eq.(14) ensures that only one among the two reactors ischosen in the Ith module. Finally, the exit from the Ithreactor module forms the inlet to any one or a combi-nation of modules I+1 to N as shown in Eq. (15). Thefeed conditions (FIf, XIf) to Ith module may be the exitconditions of any one or a combination of the previousI−1 modules (Eqs. (9) and (10) illustrate this fact).Hence, when the Ith reactor extension is considered,bypasses from the exits of the previous I−2 modulesand the feed are automatically considered. Similarly,parallel reactor structures up to the I−1 th module arealso accounted for. This structure keeps the construc-tive MINLP algorithm as compact as possible. Insidethe Ith reactor module, either the CSTR or the DSR


may be chosen. Thus, the exit conditions from the Ithreactor module (XI) are determined appropriately (Eqs.(9)–(14)). The exit stream from the Ith reactor moduleforms the inlet stream to any one or a combination ofreactor modules I+1 to Nmod (Eq. (15)). Lakshmananand Biegler (1996a) discuss the advantages and disad-vantages of this strategy. They suggest an initializationprocedure, a solution technique and remedies for possi-ble problems due to the bilinear Eqs. (10) and (13).

The Lakshmanan and Biegler (1996a) algorithm forreactor network synthesis is a versatile algorithm thatallows the designer to incorporate additional attainableregion properties when necessary. In addition, it may beeasily incorporated with techniques for simultaneoussynthesis of heat exchange networks and mass exchangenetworks (Lakshmanan & Biegler, 1996b) to simulta-neously synthesize waste minimizing process flowsheets.

4. Vinyl chloride monomer production process

The principal operating steps used in the ethylenebased balanced process for the production of vinylchloride are shown in the reactions below (Cowfer &Magistro, 1983). Ethylene is chlorinated to EDC by theprocesses of oxychlorination and direct chlorination.EDC is purified in the EDC purification section and fedto the pyrolysis reactor where vinyl chloride monomer isproduced. Vinyl chloride is purified in the vinyl chloridepurification section and the recovered hydrogen chlorideand EDC are recycled. The main reactions involved are:

Direct chlorination: C2H4+Cl12�C2H4Cl12

Oxychlorination: C2H4+2HCl+1/2O2

�C2H4Cl2+H2OEthylene dichloride 2C2H4Cl2�2C2H3Cl

pyrolysis:+2HCl

Overall reaction: 2C2H4+Cl2+1/2O2

�2C2H3Cl+H2O

The kinetics for these reactions are described belowand detailed rate expressions are given in Appendix A.Apart from the main products shown in the equationsabove, a variety of byproducts are produced in thesereactors. Some of the more significant byproducts are1,1,2 trichloroethane, chloral (trichloroacetaldehyde),cis and trans- 1,2 dichloroethylenes, mono-, di-, tri- andtetrachloromethanes and carbon oxides in the oxychlo-rination reactor; in the direct chlorination network,1,1,2 trichloroethane is also produced. Acetylene,ethylene, butadiene, methyl chloride, vinyl acetylene,chloroprene, vinylidene chloride, benzene, trichloro-ethylene, tri- and tetrachloromethane and other chlori-nated products are produced in the pyrolysis reactor.

All of these by-products cause problems and theirproduction should be minimized to improve conversionof raw materials to desired products, lower raw materialcosts, decrease the difficulties in the EDC purificationsection, prevent coking and fouling of the pyrolysisreactor and reduce waste handling and treatment costs.For example, trace quantities of impurities in the pyrol-ysis reactor feed makes the EDC pyrolysis processsusceptible to fouling and inhibition. The feed to thepyrolysis reactor should be at least 99.5% pure dryEDC. Also, the EDC recovered from the exit of thepyrolysis reactor contains an appreciable number ofimpurities, some of these like chloroprene polymerize toa rubbery material which can seriously foul the EDCpurification section. Similarly, trichloroethylene canform an azeotrope with EDC. If this byproduct is notremoved in the EDC recycle from the pyrolysis section,it can lead to reduced cracking rates and increasedfouling in this reactor. (McPherson, Starks and Fryar,1979). Hence, there is a significant need to improve theperformance of these reactors.

Fig. 2 shows a simplified process flowsheet for theproduction of vinyl chloride monomer (VCM). Thisflowsheet is based on the PRO/II Casebook (1992) onthe vinyl chloride monomer plant with a capacity of 125million lbs/year (or 56.7×106 kg/year) of VCM. Thethree reaction sections are integrated into a balancedprocess flowsheet with the separation sections (EDCpurification and VCM purification) by recycling thehydrogen chloride recovered from the VCM purificationsection to the oxychlorination section. In addition, theEDC recovered from the VCM purification section isrecycled to the EDC purification section.

Direct chlorination of ethylene to EDC is a homoge-neous catalytic reaction in the liquid phase. Commer-cially, the exothermic reaction is conducted in a liquidphase reactor by intimately mixing ethylene and chlo-rine in liquid EDC. The reaction is catalyzed by ferricchloride catalyst. The heat produced in the reaction isremoved by cooling water or by operating the reactor atthe boiling point of EDC and allowing the pure productto vaporize. Here, the latter case is considered. Wemodeled the direct chlorination reaction using kineticsderived from Wachi and Morikawa (1986) and pre-sented in Appendix A. On the other hand, in theoxychlorination process ethylene and oxygen react withdry hydrogen chloride gas produced from the pyrolysisof EDC. Commercially the reaction is conducted ineither a fixed bed (230–300°C, 1.5–14 atm gauge pres-sure) or a fluidized bed (220–235°C, 1.5–5 atm gaugepressure) reactor containing a copper chloride catalystimpregnated on a porous alumina support (Cowfer &Magistro, 1983). It is a highly exothermic reaction andgood temperature control is essential. Increasing reac-tor temperature results in increased byproduct forma-


Fig. 2. Vinyl chloride monomer production process flowsheet.

tion and catalyst deactivation. The reaction is mod-eled using the kinetics given in Appendix A (obtainedfrom a series of papers by Gel’perin, Bakshi, Avetisov& Gel’bshtein, 1979; Gel’perin, Bakshi, Avetisov &Gel’bshtein, 1983; Gel’perin, Bakshi, Avetisov &Gel’bshtein, 1984).

Trace quantities of impurities in the pyrolysis reac-tor feed can cause fouling of the reactor. Hence,EDC from three different sources: direct chlorination,oxychlorination and EDC recovered and recycledfrom the pyrolysis section, must be purified to atleast 99 wt.% pure EDC. The EDC produced in theoxychlorination section is washed in caustic to re-move unreacted hydrogen chloride. This stream ismixed with the EDC produced in the direct chlorina-tion section and fed to the EDC purification section.The EDC purification section consists of a decanterfollowed by two distillation columns. The first columnis the lights column where 99.9% of the water is re-moved as top product. The bottom product of thiscolumn is fed into the heavies column. The topproduct of the heavies column is 99.9 wt.% pureEDC which is fed to the pyrolysis section, as ex-plained in the next paragraph. The VCM produced inthe pyrolysis section is separated from the HCl andEDC in the VCM purification section. The firstcolumn in the VCM purification section, called theHCl column, receives mixed phase feeds containing

mainly EDC, VCM and HCl with small amounts ofreaction byproducts. HCl distills off at the top; halfof it is recycled to the oxychlorination section andthe rest is refluxed back to the column. The bottomproduct is fed to the VCM column, where purifiedVCM is separated out as the overhead product andthe recovered EDC is recycled to the EDC purifica-tion section.

Finally, vinyl chloride monomer is produced bythermal cracking of EDC. The endothermic reactionis carried out commercially in tubular reactors attemperatures of 480–530°C and reactor gauge pres-sures of 6–35 atm (Ranzi, Grottoli, Bussani Che &Zahng, 1993). The main reaction which yields vinylchloride monomer and hydrogen chloride is a homo-geneous, first-order free-radical chain mechanism. Sev-eral byproducts are produced in trace amountsthrough related free-radical and molecular mecha-nisms. The EDC pyrolysis process is very selective tovinyl chloride but it is difficult to verify the forma-tion and amount of the byproducts due to propri-etary kinetics. Also, literature models for the kineticparameters of the reactions involved are not compre-hensive. The pyrolysis reaction kinetics used in thisstudy were compiled from various literature sources(Kurtz, 1972; Weissman & Benson, 1984; Karra &Senkan, 1988; Ranzi, Dente, Tiziano, Mullick & Bus-sani, 1990) and described further in Appendix A.


5. Results of the vinyl chloride case study

Eq. (P1) as well as the flowsheet connectivity andseparators shown in Fig. 2 were modeled in GAMS(Brooke, Kendrick & Meeraus, 1988). The resultingmodel is a MINLP (which we refer to as P1f) with4236 constraints, 4995 continuous variables and eightbinary variables. Four binaries were used in each ofthe direct chlorination and oxychlorination sections tosynthesize the optimal reactor structures. Based onliterature reports, a tubular reactor (PFR) in a fuelfired furnace was used to model the pyrolysis reactor.The purification sections are modeled using split frac-tions, which were based on the data published in thePRO/II Casebook (1992). Each of the reactors in themodel was initialized with profiles obtained fromsolving segregated flow initial value problems. TheMINLP model was solved using DICOPT(Viswanathan & Grossmann, 1990).

5.1. Vinyl chloride yield maximization

We first consider the flowsheet where we maximizethe overall conversion of ethylene to vinyl chloridemonomer. Model (P1f) was solved in 8.12 s on aHP-UX 9000-720 workstation. The results of thisflowsheet optimization are shown in Tables 1 and 2.In addition we can obtain trade-off curves betweenthe conversion of ethylene and the amount of wasteallowed by solving the following sub-problems:

Primary Problems (P2)Max VCM Min W(x, y)

Yield (x, y)x, y x, y

h(x, y)=0 h(x, y)=0g(x, y)50g(x, y)50

W(x, y) W(x, y)

5Wallowed 5Wallowed

x�X x�X

¤Rm y�Y¤Rm y�Y

={0, l}q ={0, l}q

Secondary ProblemsMin VCM yieldx, y

h(x, y)=0g(x, y)=0W(x, y)=o

x�X�Rm y�Y�Rq

o� [Wmin, Wallowed]

where x is the set of continuous flowsheet parameters,y is the set of binary variables and h(x, y) andg(x, y) are the constraints from (P1f) and o is aparameter varied between Wmin and Wallowed. SeeLakshmanan and Biegler, (1995) for more details.From the solutions of these MINLP models, we nextdescribe the results for each reaction section for theyield maximization case.

Table 1Results of the vinyl chloride case study

PurificationChlorination

Direct Oxy- Pyrolysis EDC VCM

Raw material*0.533 0.481C2H4 7.0e−6 7.0e−3 7.0e−6

1.3e−5l.5e−31.6e−6Cl2 0.5281.6e−40.241O2

Intermediate*EDC 0.521 0.480 1.67 1.0e−3 0.67

By-product*5.798e−3 5.582e−4C2H3Cl3 6.3e−3 1.6e−5

CO2 1.622e−4 1.6e−4C2H2,

2.9e−5C4H6,C2H2Cl2

Main productC2H3Cl 1

* Mol/mol of vinyl chloride product.


Table 2Reactor conditions that maximize conversion of ethylene to vinyl chloride

Reactor feed (gmol/s)Reactor conditions (residence time, pressure, temperature, conversion)Reactor struc-ture

CSTR 9293 s, 1 atm, 333 K, 99.7% 16.00 C2H4 15.84 Cl2Direct chlorina-tion

324 s, 5.3 atm, 466.3–465.8 K, 96.2% 14.43 C2H4 7.26 O2Oxy-chlorination PFR+5.6 s, 5.3 atm, 465.7 K, 96.24% 30.03 HClCSTR0.35 s, 20 atm, 773 K, 60% 49.99Pyrolysis PFR

5.2. Direct chlorination of ethylene

In this scheme (Wachi & Morikawa, 1986, see Ap-pendix A), EDC is formed by an addition reactionbetween ethylene and chlorine. The main byproduct is1,1,2 trichloroethane which is formed by substitutionand addition reactions between chloride radicals, chlo-rine and EDC. The temperature range considered was30–75°C, the pressure range was 0–2 atm. A slightexcess of ethylene was used. The optimal reactor struc-ture determined by using the reactor network synthesisalgorithm (Eq. (P1)) was a CSTR (9293.48 s residencetime). This yielded a 99.70% conversion of the limitingcomponent (chlorine) and 99% selectivity to EDC.

5.3. Oxychlorination of ethylene

Here EDC is formed through a heterogeneous cata-lytic cycle where the main chlorinating agent is cop-per(II) chloride. The byproducts considered aretrichloroethane and carbon oxides. The recycled hydro-gen chloride stream from the top of the vinyl chloridemonomer purification section is the limiting componentin this reaction. The optimal reactor network deter-mined by using the reactor network synthesis al-gorithm (P1) was a PFR with a residence time of 324 sand 96.20% conversion of the rate limiting component(hydrogen chloride), followed by a CSTR with a resi-dence time of 5.6 s and 96.24% conversion of thelimiting component, with a selectivity of ethylene toEDC of 99.8%. Note that the small CSTR is neededfor the additional conversion, optimization with thePFR alone leads to a conversion of only 96.205%. Theoptimal temperature profile in the PFR is shown inFig. 3 and the reactor pressure was found to be 5.34atm.

To compare our reactor network to that used in thevinyl chloride monomer plant PRO/II (1992) casestudy, we consider their reactor network, a PFR at290°C and 6.122 atm, modeled using the kineticscheme in Appendix A. The results for this base casereactor showed that the amount of trichloroethaneproduced per kg of EDC formed is 0.016 kg and theamount of carbon oxides produced per kg of EDCformed is 0.03 kg, as compared to our results of

0.00154 and 0.00015 kg, respectively, using the reactornetwork (PFR+CSTR) and conditions shown above.Hence, a waste reduction of one and two orders ofmagnitudes is realized.

Also, using MINLP (P2) we can plot the noninferiorcurve, shown in Fig. 4. between conversion of ethyleneto EDC and the amount trichloroethane produced inthe oxychlorination reactor. The optimal reactor struc-tures for this curve remained a PFR+CSTR combina-tion for the conversion maximization case.

5.4. Pyrolysis of EDC

Vinyl chloride monomer is produced by thermalcracking of EDC and this endothermic reaction iscarried out commercially in tubular reactors at temper-atures of 480–530°C and gauge pressures of 6–35 atm(Ranzi et al., 1993). The byproducts considered in thekinetic model include acetylene, ethylene, butadiene,trichloroethane and vinyl acetylene. Each of thesebyproducts are harmful and could foul the reactor, ifrecycled. A plug flow reactor was used to model thepyrolysis section using the kinetics in Appendix A.This choice was based on solutions obtained fromsolving initial value problems, which show very highselectivity to vinyl chloride in a plug flow reactor and

Fig. 3. Temperature profile in the PFR-oxychlorination section.


Fig. 4. Noninferior curve between ethylene conversion and byproductformation in the oxychlorination process.

Table 4Prices of utilities

Utility Cost

Cooling water (298–319 K) 1.865e−5 $/kg7.694e−3 $/kg300 psig steam

Steam CondensateFuel Oil 1.9162e−8 $/JRefrigerant 5e−8 $/J

structures and conditions that maximize the conversionof ethylene to vinyl chloride. The overall conversion ofethylene to vinyl chloride is 98.7% (see Table 1;1/(0.533+0.481)).

So far we have assessed the performance of thereactor networks based on conversion to the desiredproduct. In addition, we synthesized maximum yieldreactor networks for a given waste limit. This allowedus to establish an optimal trade-off curve of conversionof raw material versus waste generated, as shown inFig. 4. However, an economic optimization of thisprocess could be of more interest to the designer. Thenext section presents a preliminary cost analysis of theproposed process flowsheet and highlights changes inreactor structures when the objective is to maximizeprofit. This analysis is then extended to consider energyintegration of the flowsheet.

5.5. Economic optimization of the 6inyl chloridemonomer production process

The aim of this study is to analyze changes in reactorstructures when the objective is to maximize pre-taxprofit instead of conversion. The vinyl chloridemonomer flowsheet shown in Fig. 2 is used in thisstudy. The objective function, the pre-tax profit, in-cludes the product and raw materials as well as utilityand equipment costs. The substitution of the objectivefunction in Eq. (P1) forms the MINLP (P1e).

The prices of the product and raw materials, heatcapacity data and utility prices are given in Tables 3and 4. The expressions used for the installed costs ofthe reactors and columns are given in Table 5 in $/yraccording to data from Turkay and Grossmann (1996).The installed costs of the reactors are based on the

from literature reports on the EDC pyrolysis process.The temperature and pressure of the PFR were 500°Cand 20 atm, respectively. The selectivity of EDC tovinyl chloride in this PFR was found to be 99.9% witha single pass conversion of 60%.

The results of this study are summarized in Tables 1and 2. In Table 1, the values shown in the intermediatessection represent the composition of products of oxy-chlorination and direct chlorination and the composi-tion at the inlet to the pyrolysis reactor (moles ofcomponent/mole of vinyl chloride produced). In thebyproducts section, the values shown are the composi-tions of these byproducts produced in the reactors. Thecompositions of acetylene, butadiene and vinylidenechloride produced in the pyrolysis reactor are summedup in the last column. The values listed under thepurification sections represent compositions of streamsleaving the process. The results shown in Table 1 matchpublished industrial data (McPherson et al., 1979) withalmost equal amounts of EDC produced in the oxy-chlorination and direct chlorination sections. This con-forms to stoichiometric requirements since the processflowsheet considered in this study is a balanced produc-tion process flowsheet and for every mole of EDCpyrolyzed to vinyl chloride monomer, a mole of HCl isformed. This HCl is completely consumed in the oxy-chlorination reactor. Table 2 gives the optimal reactor

Table 3Prices and heat capacity data of the products and raw materialsa

Cl2 HClC2H4 O2 C2H4Cl2 H2OC2H3Cl

Prices @ $/kg 0.5346 0.46240.1928 1.5891 0.0275 0.3749−9.236e405.234e4 0DHf(298) J/gmol 3.517e4 −1.298e5 −2.42e530.67 32.2426.933.806Cpvap J/gmol-K 20.495.94928.11

a Chemical and marketing reporter, 12, 1996. Reid, Prausnitz and Poling, 1986. DHf(T)=DHf,298.2+Cp (T−298.2); DHr×n=SDHf(products)–SDHf(reactants).


reactor volumes (flowrate×residence time) and fixedvalues of g(P) and d(T) were assumed for each of thereactors depending upon their temperature and pressure.The installed costs of the column are determined basedon the throughput through the column. In Fig. 2, thedirect chlorination and oxychlorination reactors areexothermic, while the pyrolysis reaction is endothermic.The heat produced in the direct chlorination reactors isremoved by operating the reactor close to the boilingpoint of EDC and allowing the pure product to vaporize.The heat loads of the rest of the process are supplied bythe utilities shown in Table 4. Hence, the heat producedin the oxychlorination reactor is removed by circulatingcold water and the effluent stream from the reactor iscooled in the oxychlorination aftercooler before enteringthe EDC purification section. In the EDC purificationsection the heat required by the reboilers is supplied byfuel oil in the lights column and 300 psig steam in theheavies column and cold water is used to cool thecondensers. A constant temperature difference is as-sumed between the reboiler and the condenser (An-drecovich & Westerberg, 1985). The reflux ratios shownin Fig. 2 are control specifications. The top product ofthe heavies column, which is about 99.5% pure dry EDC,is compressed and heated to 500°C (by using 300 psigsteam and fuel oil) in the pyrolysis preheater. The heatrequired by the endothermic pyrolysis reaction is sup-plied by fuel oil and the exit stream from the pyrolysisreactor is cooled by cooling water in the after-cooler.Finally, the HCl column recovers hydrogen chloridewhich is recycled to the oxychlorination section and theVCM column produces pure vinyl chloride monomer asthe top product. The condensers in both columns requirea refrigerant while 300 psig steam is used to supply heatto the reboilers.

The optimal reactor structures from the solution of(P1e) were found to be: a CSTR for direct chlorination,a PFR for oxychlorination and a PFR for the pyrolysissection. The reactor conditions are similar to thoseobtained for the conversion maximization case. Thisshows that the improvement in conversion which may beachieved by using the additional reactor structure (thesmall CSTR) described in the previous section is notenough to offset the increased costs incurred by using

Table 6Summary of economic optimization

No heat integration With heat integration

Reactor structure(res. time s, press.atm, temp. K)

CSTR(9293, 1, 333)Direct chlorination CSTR(9427, 1, 333)

PFR(324, 5.2, 466.3–465.7) PFR(324, 5.l, 466.3–465.7)OxychlorinationPFR(0.35, 20, 773)Pyrolysis PFR(0.35, 20, 773)

0.01116 0.01122Reactor installedcosts 106 $/year

0.00456 0.00456Column installedcosts 106 $/year

Product & raw ma-terial flows kg-mol/year

C2H3Cl 907.135 907.135C2H4 919.786919.862

463.772Cl2 463.697226.281O2 226.295

98.63%O6erall con6ersion 98.62%of C2H4 toC2H3Cl

Process utility costs106 $/year

0.8603300 psig steam 0.39750.69950.8852Fuel oil

Cooling water 0.05080.0856Refrigerant 1.47771.7423

0.9479Savings due to heatintegration

3.2762.327Profit* 106 $/year

* Profit=Product value—raw material costs—reactor and columninstalled costs–process utility costs.

them. As a result, the reactor networks are simplified.The pre-tax profit realized by this flowsheet was $2.33million/year (based on 8400 h of operation/year). Theoptimization problem contained 4265 variables, eightbinary variables and 5024 constraints; it was solved in 5.6CPU s on a HPUX 9000-720 from a good initial point.Results of this problem are summarized in Table 6 alongwith reactor designs and operating conditions.

5.6. Sensiti6ity of the 6inyl chloride monomerproduction process

In this section, we take a brief look at the sensitivityof the VCM flowsheet with respect to changes in thekinetic parameters in the direct chlorination and oxy-chlorination reaction sections. In particular, we want toshow that even with reasonable changes in the values ofthe kinetic constants, overall conversion to VCM re-mains quite high with the reported optimal reactornetwork and that the annual process profit is not

Table 5Installed costs of the reactors, columns and heat exchangersa

Reactor (MS/280) Fm 0.125 VR g(P) d(T)1000 $/year

Column 0.1512 (throughput)0.5 1000 $/year1000+560[Area (m2)]0.6 $/yearHeat exchangers, heaters

and coolers

a Turkay and Grossmann, 1996. MS=1031.8 for 1996; Fm=2.25for SS; VR=reactor volume; g(P)=1 for PB3.4 MPa; d(T)=1 forT=300–400 K, d(T)=1.1 for T=400–500 K, d(T)=1.4 for T=700–1000 K.


significantly affected. In addition, if the flowsheet isreoptimized according to the new kinetic values, thereactor structures in each section of the plant do notchange significantly.

We investigated six scenarios for the direct chlorinationsection in which each rate constant in reactions DC1–DC3 (see Appendix A) is allowed to change from itsnominal value by 910% while the other two were heldfixed. The flowsheet was simulated using the reactorstructures found in Table 2 with the new values of therate constants and the change in overall conversion ofethylene to VCM was noted. For all cases, the overallconversion to VCM changed less than one percent withthe largest decrease in conversion occurring when rateconstant for reaction DC1 was 10% lower than its nominalvalue. If we consider an unfavorable change in the rateconstants with a 20% decrease in k1 in reaction DC1 (lowerrate of ethylene dichloride formation) coupled with a 20%increase in rate constants in reactions DC2 and DC3(increased byproduct formation), the overall conversionto VCM is still quite high for the reactor structure in Table2 with a slightly lower conversion of 98.1%.

For the oxychlorination section, the pre-exponentialfactors for reactions O1–O3 were changed by910%. Theequilibrium constant for reaction O4 was also changedby 910%. With these variations, the overall conversionchanged less than 1% when using the reactor structurefrom Table 2, with the largest decrease in conversionrealized when the rate constant in reaction O1 was 10%lower than its nominal value. Even with a 20% decreasein the rate constant for reaction O1 and a 20% increasein the rate constants for reactions O2–O3 (we assume anominal value of the equilibrium constant in reaction O4here), the overall conversion to VCM decreases less thanone percent to 98.2%. Thus, with small changes in thedirect chlorination and oxychlorination rate constants,overall conversion is quite good for the optimal reactornetwork found in Table 2.

We also reoptimized the flowsheet for the new valuesof the kinetic parameters to see what changes would berealized in the reactor networks. All fresh feeds to theprocess were fixed to the values reported in Table 2. Forthe oxychlorination section, the inlet temperature wasfixed to 466.3 K. For each case, the flowsheet wasinitialized using the results from Table 2. In all of thereoptimized flowsheets, the reactor structures remain thesame as reported in Table 2 with minor changes to thereactor residence times.

With varying rate constants in the direct chlorination,the residence times of the reactors in both the directchlorination section and oxychlorination section changedfor maximizing overall conversion to VCM. For the rateconstants in reactions DC2 and DC3, very little effect wasseen on the direct chlorination CSTR residence time. Inthe worst case for these two reactions, the optimalresidence time increased or decreased by less than half

a percent, clearly a negligible change. In the oxychlorina-tion, both the PFR and CSTR also changed less than ahalf percent from the values reported in Table 2. Witha decrease of 10% in the rate constant for reaction DC1,the optimal residence time in the CSTR increased to10 314 s. This is reasonable since reaction DC1 producesthe desired product, EDC. The oxychlorination PFRdecreased to 292 s and the CSTR decreased slightly to5.1 s. This new structure maintains a very high overallconversion of ethylene to VCM of 98.6%. With a 10%increase in the rate constant for reaction DC1, the optimaldirect chlorination CSTR residence time decreased to8457 s while the oxychlorination PFR increased to 392s followed by a CSTR with a residence time of 38 s. Overallconversion to VCM for this scenario is 98.8%.

For the oxychlorination section, eight optimizationproblems were solved corresponding to a 910% changein each of the four kinetic parameters. The residence timeof the oxychlorination PFR increased to 396 s when therate constant for reaction O1 was 10% lower than itsnominal value. The oxychlorination CSTR also increasedslightly to 41.1 s while the CSTR in the direct chlorinationdecreased to 9183 s. Overall conversion of ethylene toVCM was 98.5%. For all other combinations, the size ofthe oxychlorination reactors didn’t change by more than2% from the values reported in Table 2. The largestincrease in residence time for the direct chlorinationCSTR was 8.8% to 10 115 s when the rate constant infor reaction O3 was 10% larger than its nominal value.For all other instances of varying oxychlorination kineticparameters, the direct chlorination CSTR changed lessthan 3% from the value reported in Table 2.

For the economic optimization (without heat integra-tion) of the VCM plant, the annual profit changes onlyslightly with variations in the direct chlorination andoxychlorination kinetic parameters. If the reactor net-works are fixed to those reported in Table 6, the overallprofit decreases for all scenarios of the varying kineticparameters. The largest decrease in profit was 5.1% to2.208*106 $/year when the rate constant in reaction DC1was 10% lower than its nominal value. If the flowsheetis reoptimized for this worst case, the new optimalstructure contains a CSTR for the direct chlorination(10 314 s), a PFR in the pyrolysis section (0.35 s) and PFR(292 s) in the oxychlorination section with nearly identicaloperating conditions as reported for the nominal case. Theannual profit is 2.267*106 $/year.

These results show that the optimal reactor networksfound for ethylene conversion to VCM and annualprocess profit are only mildly sensitive to changes in thekinetic parameters.

5.7. Energy integration of the process flowsheet

Obviously, the heat content of the reactor streams andhot streams in the aftercoolers and condensers may be


used to heat cold streams in the preheaters andreboilers. Eight hot process streams (H1…H8) and ninecold process streams (C1…C9) are shown in Fig. 2. Asmentioned earlier, all the heat produced in the directchlorination reactor is used to vaporize the liquid EDCproduct. Hence, the streams flowing through this reactorare not considered while energy integrating the processflowsheet. The oxychlorination reactor is exothermic.The heat produced here may be used to heat otherprocess streams. Since the temperature drop across theoxychlorination reactor is small, the reactor is repre-sented by two hot streams (H1 and H2) (instead of twohot streams per finite reacting element). Hot stream H1accounts for the heat of reaction (a 1 K temperature dropis assumed) and H2 accounts for the hot reactor streamswith a heat capacity equal to that of the reactor internalstreams. The inlet temperature of streams H1 and H2 isthe temperature at the inlet to the reactor and the exittemperature of stream H2 is equal to the temperature atthe reactor exit. The temperature at the reactor exit isoptimized simultaneously with the rest of the process andenergy network constraints. Similarly, the endothermicpyrolysis reactor is represented by two cold streams C6and C7, one of which (C6) accounts for the heat ofreaction and the other stream (C7) may be heat inte-grated with the rest of the process. Unfortunately, thetemperature at which the pyrolysis reactor is operated,773 K, is much higher than the temperature of the otherprocess streams. Hence, these streams require an externalheating utility. Similarly, the lights column reboilerstream (cold stream C2) and the HC1 column condenserstream (hot stream H7) require an external heating andcooling utility, respectively. The rest of the streams maybe energy integrated.

The simultaneous energy integration technique pro-posed by Duran and Grossmann (1986) is used todetermine the minimum utility consumption of thisprocess flowsheet. This requires the addition of thefollowing equations to MINLP formulation (P1e):

QC=QH+ShoH(FCp)h[Thin−Th

out]

−ScoC(FCp)c[T cout− t c

in]

QH]zHp (x, y)

Here, (FCp) is the heat capacity-flowrate and Tp corre-sponds to the pinch candidates which are derived fromTh

in for the hot streams and t cin+DTm for the cold

streams. Also, we have

zHp (y)=%coC(FCp)c

× [max {0; t cout−{Tp−DTm}}

−max {0; t cin−{Tp−DTm}}]

−%hoH(FCp)h[max {0; Thin−Tp}−max {0; Th

out−Tp}]

with the max operator approximated with a smoothingfunction (see Balakrishna & Biegler, 1993). We augment(P1e) with these equations to form the MINLP (P1h). Aminimum approach temperature (DTm) of 10 K is as-sumed for efficient heat exchange. For this energyintegrated case, the pre-tax profit for the same vinylchloride production rate was found to be 3.27 million$/year. This is mainly due to savings in the utility costsand slightly better conversions in the reactors. The resultsof the optimization (4432 variables, eight binary vari-ables, 5168 constraints, 6.96 CPU s on a HP-UX9000-720) are summarized in Table 6. Note that heatintegration of the process flowsheet does not affect theoverall conversion significantly because the overall con-version is high even without heat integration. Also, as thisis a difficult problem to solve computationally, thesolution times shown above are for a problem with agood initialization.

Following the solution of (P1h), the resulting temper-atures and heat capacity flowrates are used to synthesizea heat exchange network (HEN). This is determinedusing the MINLP formulation proposed by Yee andGrossmann (1990). The SYNHEAT2 program devel-oped by Bolio (1994) to implement this MINLP wasmodified to solve this large HEN problem. The modifi-cations include, relaxing the upper bound in the calcula-tion of approach temperatures and a user-supplied initialset of stream matches. The MINLP model for thesynthesis of the HEN involved 16 stages (7922 equalitiesand inequalities, 7394 continuous variables). The heatcapacity data for the individual streams is given in Table7. The intermediate utilities (cold water, steam and steamcondensate) were considered as process streams. 9 hotstreams (H1–H6, H8, steam and steam condensate)and 7 cold streams (C1, C3–C5, C8, C9, cold water)

Table 7Data for the HEN synthesis in the vinyl chloride case studya

Stream Tin (K) Tout (K) F (KW K-1)

466.32 465.32H1 849.71466.32H2 465.77 1.54

H3 465.77 330.50 1.37H4 608.14394.07 393.07

420.16 419.16H5 1600.24330.62772.79H6 6.46

H7 484.48240.80241.80309.26H8 308.26 667.81298.20C1 466.32 1.19481.07C2 482.07 12.32

C3 431.66 432.66 2560.38419.16C4 473.00 0.05

772.00 0.06473.00C5773.00772.00 1206.47C6

772.00 772.79 0.13C7C8 2699.90359.80 360.80C9 430.47 1392.21431.47

a h=1 kW m−2 K−1 for all the streams. Cost of Heat Exchangers,Heaters and Coolers ($ year−l)=1000+560(Area(m2))0.6.


Fig. 5. Heat exchange network for the process streams.

were included in the MINLP model. Hot stream H7 wasmatched with the refrigerant and cold streams C2, C6and C7 were matched with heat from fuel oil. Thenetwork shown in Fig. 5 indicates heat exchange betweenthe following streams labeled in Fig. 2.

As seen in Fig. 5, the HEN includes the followingmatches:� oxychlorination reactor streams (H1) and the VCM

column reboiler stream (C9);� oxychlorination reactor streams (H2) and the oxy-

chlorination preheater stream (C1);� heavies column condenser stream (H5) and the HCl

column reboiler stream (C8);� pyrolysis after-cooler stream (H6) and the pyrolysis

preheater streams (C4, C5);

� pyrolysis after-cooler stream (H6) and the heaviescolumn reboiler stream (C3); and

� pyrolysis after-cooler stream (H6) and the VCMreboiler stream (C9);

This network consumes about 16 KW more of theheating and cooling utility as compared to the mini-mum utility target found by (P1e). The capital costs ofthis network were about $70 000/year (calculated bythe cost expressions shown in Table 5), which is consid-erably smaller than the savings realized by the energyintegration.

6. Conclusions

The reactor network synthesis algorithm was applied


for the improvement of the vinyl chloride monomerproduction process. The approach has been successfullymodeled for the three reaction sections (oxychlorination,direct chlorination and pyrolysis) and the conversion ofethylene to vinyl chloride has been maximized. For thiscase, the three reaction sections were found to be:Oxychlorination: PFR+CSTR (small); Direct chlorina-tion: CSTR; and a PFR for the pyrolysis section.Compared with the results of an industrial case study,these reactor structures perform better than the conven-tional reactors used for these processes and enhance theconversion of raw materials to the desired products whileminimizing waste production. Also, the noninferiorcurve between one of the main waste products and theconversion of ethylene was plotted and a simple processflowsheet which incorporates these results has beenproposed.

Moreover, an economic optimization of the proposedprocess flowsheet was presented and changes in reactorstructures while optimizing the pre-tax profit were high-lighted. A simultaneous energy integration of the processflowsheet was also performed to determine the minimumutility consumption of the process and a heat exchangenetwork was derived. The network indicated efficient useof the oxychlorination reactor and pyrolysis aftercoolerstreams to heat other process streams. This results insignificant cost savings. Future research will continue toaddress the sensitivity of reactor networks to process andkinetic parameters used in the reaction schemes.

7. Nomenclature

R(X(a), T(a)) residence time distribution in the PFRf(aI) residence time distribution in the PFR

or DSR in module Iinlet volumetric flowrate to the Ith re-Fif

actor moduleFIc, FId volumetric flowrate through the CSTR

and DSR in the Ith reactor moduleFkI−1 exit volumetric flowrate from module

k which is an inlet to module I−1,k=0, I−1

Ne the number of finite reacting elementsin the reactor discretization

Nmod the number of reactor modules consid-ered in the synthesis of the reactornetworkfraction of Qside entering the DSR atq(aI)a in module Ivolumetric flowrate through the PFRQ(aI)or DSR at a in reactor module I

Qside side stream flowrate to the DSRR(X(a), T(a)) rate vector at a

rate vector at element i, collocationR(Xij, Tij)point j

residence timet

temperature inside the reactor at aT(a)temperature inside the CSTR in mod-TCSTR I

ule ITDSR exit I temperature at exit of the DSR in

module IU the upper bound on the variable

mass concentration vector at point aX(aI)inside the DSR in module Imass concentration vector inside theXCSTR I

CSTR in module Imass concentration vector inside theXDSR I

DSR in module IXicin mass concentration vector at the inlet

to the CSTR in module Imass concentration vector at the inletXIdin

to the DSR in module Imass concentration vector at the feedXif

to module Ibinary variable denoting the existence/Yic

nonexistence of the CSTR in module IYid binary variable denoting the existence/

nonexistence of the DSR in module I

Acknowledgements

Funding from the Engineering Design Research Cen-ter, an NSF sponsored Engineering Research Centerand from the Pennsylvania Infrastructure TechnologyAlliance (PITA) is gratefully acknowledged.

Appendix A. Kinetic schemes for the three reaction sectionsin the vinyl chloride monomer production process.

A.1. Direct chlorination reaction kinetics

The kinetics for direct chlorination of ethylene toEDC (at 60°C) were derived from (Wachi & Morikawa,1986). EDC and trichloroethane are formed throughmolecular and radical addition and substitutionreactions.

Reaction DC1:

C2H4+Cl2�k1

C2H4Cl2

R1=k1[C2H4][Cl2]

k1=0.132 m3 mol−1 s−1

Reaction DC2:

C2H4+Cl2�k2

C2H4−Cl+Cl�

Cl�+C2H4Cl2�k3

C2H3Cl2�+HCl

C2H3Cl2�+Cl2�k4

C2H3Cl3+Cl�

C2H3Cl2�+Mi�k5

deactivate


Cl�+Mi�k6

deactivate

R2=kR2[C2H4][Cl2]2

kR2=0.0239 m6 mol−2 s−1

Reaction DC3:

C2H4Cl2+Cl2�kR 3

C2H3Cl3+HCl

R3=kR3[C2H4Cl2][Cl2]

kR3=6.12�10−9 m3 mol−1s−1

A.2. Oxychlorination reaction kinetics

The kinetics for oxychlorination of ethylene to EDCwere derived from a series of papers by (Gel’Perin etal., 1979, 1983, 1984). EDC and several byproducts areformed through molecular and radical addition andsubstitution reactions. The main reactions are shownbelow.

Reaction O1:

C2H4+2HCl+1/2O2�k1

C2H4Cl2+H2O

k1=a1* exp(b1/RT) mol/(liter cat. h kPa1.5)

a1=104.2

b1= −40.1 kJ R=8.314 J/mol K

rC2H4=k1pC2H4

pCl20.5 mol/(liter cat. h)

Reaction O2:

C2H4Cl2+HCl+1/2O2�k2

C2H3Cl3+H2O


a2=1013.23 b2= −128.04 kJ

rC2H4Cl2=k2pC2H4Cl2

pCl20.5 mol/(liter cat. h)

Reaction O3:

C2H4+3O2�k3

2CO2+2H2O


a3=106.78 b2= −112 kJ

rC2H4

ox =k3pO2pCl2

0.5pC2H4mol/(liter cat. h)

Reaction O4:4HCl+O2l

K4

2Cl2+2H2O

K4=k4f/krb

k4f=1000�exp(17.13−13000/(1.987�T))

k4b=exp(5.4+16000/(1.987�T))

rCl2=K4pO2pCl2

−1 mol/(liter cat. h)

A.3. Pyrolysis reaction kinetics

The kinetics for the pyrolysis of EDC to vinyl chlo-ride were derived from several papers (Kurtz, 1972;Weissman & Benson, 1984; Karra & Senkan, 1988;Ranzi et al., 1990). Vinyl chloride and several byprod-ucts are formed through molecular and radical additionand substitution reactions. The main reactions whichfollow elementary reaction kinetics are shown below.

Reaction P1:

1,2C2H4Cl2�K1

C2H3Cl+HCl

K1=k1f/k1b

k1f=1013.6 exp(−58000/(1.987�T)) s−1

k1b=0.3�109 exp(−44000/(1.987�T)) m3 kmol−1 s−1

Reaction P2:

C2H3CllK2

C2H2+HCl

K2=k2f/k2b

k2f=0.5�1014 exp(−69000/(1.987�T)) s−1

k2b=0.37�109 exp(−40000/(1.987�T)) m3 kmol−1 s−1

Reaction P3:

1,2C2H4Cl2�K3

C2H4+Cl2

k3=1013 exp(−72000/(1.987�T)) s−1

Reaction P4:

C2H4lK4

C2H2+H2

K4=k4f/k4b

k4f=0.1�1015 exp(82000/(1.987�T)) s−1

k4b=0.8�109 exp(−38000/(1.987�T)) m3 kmol−1 s−1

Reaction P5:

C2H2+C2H4lK5

C4H6

K5=k5f/k5b

k5f=0.15�109 exp(−32000/(1.987�T)) m3 kmol−1 s−1

k5b=0.5�1013 exp(−73000/(1.987�T)) s−1

Reaction P6:

C2H3Cl+CllK6

CH2Cl CH Cl

K6=k6f/k6b

k6f=0.2�1011 exp(−0/(1.987�T)) m3 kmol−1 s−1

k6b=0.15�1014 exp(−19000/(1.987�T)) s−1

Reaction P7:

1,2C2H4Cl2�k7

CH2 Cl CH Cl+HCl

k7=1010.8 exp(−2900/(1.987�T)) m3 kmol−1 s−1


Reaction P8:

1,2C2H4Cl2+Cl�K8

CH2 CH2 Cl+Cl

k8=0.2�1017 exp(−83000/(1.987�T)) m3 kmol−1 s−1

Reaction P9:

C2H3Cl+ third�k9

C2H3+Cl+ third

k9=0.25�1016 exp(−78000/(1.987�T)) s−1

Reaction P10:

CH2CH2Cl�k10

C2H4+Cl

k10=0.1�1015 exp(−19000/(1.987�T)) s−1

Reaction P11:

C2H3+ third�k11

C2H2+H+ third

k11=0.8�1012 exp(−30600/(1.987�T)) s−1

Reaction P12:

C2H3Cl+H�k12

CH2CH2Cl

k12=0.1�1011 exp(−0/(1.987�T)) m3 kmol−1 s−1

Reaction P13:

C2H3+Cl+ third�k13

C2H3Cl+ third

k13=0.1�1011 exp(+13000/(1.987�T)) m3 kmol−1 s−1

Reaction P14:

C2H3+C4H6�k14

C6H9

k14=108.5 exp(−3000/(1.987�T)) m3 kmol−1 s−1

Reaction P15:

C6H9�k15

C6H8+H

k15=1013.5 exp(−42200/(1.987�T)) s−1

Reaction P16:

C6H8�k16

C6H6+H2

k16=1013.5 exp(−40000/(1.987�T)) s−1

Reaction P17:

CH2ClCHCl+Cl2�k17

Cl+CH2ClCHCl2

k17=108.8 exp(−1000/(1.987�T)) m3 kmol−1 s−1

Reaction P18:

CH2ClCHCl2lK18

C2H2Cl2+HClK18=k18f/k18b

k18f=0.2�1014 exp(−58000/(1.987�T)) s−1

k18b=0.3�109 exp(−44000/(1.987�T)) m3 kmol−1 s−1

Reactions P9, P11, P12, P13, P14, P15 and P16 werenot used in the GAMS model for the pyrolysis reactor,

since only trace amounts of these radicals were pro-duced and they caused numerical inconsistencies in themodel.

References

Andrecovich, M., & Westerberg, A. W. (1985). An MILP formula-tion for heat integrated distillation sequence synthesis. AIChE J.,31, 363.

Aris, R. (1960). Studies in optimization. I: the optimum design ofadiabatic reactors with several beds. Chemical Engineering Sci-ence, 12, 243.

Aris, R. (1961). The optimal design of chemical reactors. New York:Academic Press.

Balakrishna, S., & Biegler, L. T. (1992a). A constructive targetingapproach for the synthesis of isothermal reactor networks. I & ECResearch, 31(9), 300.

Balakrishna, S., & Biegler, L. T. (1992b). Targeting strategies forsynthesis and energy integration of non-isothermal reactor net-works. I &EC Research, 31(9), 2152.

Balakrishna, S., & Biegler, L. T. (1993). A unified approach for thesimultaneous synthesis of reaction, energy and separation systems.Industrial and Engineering Chemistry, Research, 32(7), 1372.

Bolio, B. (1994). Guidelines for the SYNHEAT interface. CarnegieMellon University: Department of Chemical Engineering.

Brooke, A., Kendrick, D., & Meeraus, A. (1988). GAMS: a user ’sguide. Redwood City, CA: Scientific Press.

Chitra, S. P., & Govind, R. (1981). Yield optimization for complexreactor systems. Chemical Engineering Science, 36, 1219.

Chitra, S. P., & Govind, R. (1985). Synthesis of optimal reactorstructures for homogenous reactions. AIChE Journal, 31(2), 177.

Cowfer, J. A., & Magistro, A. J. (1983). Vinyl chloride and poly(vinylchloride). Kirk Othmer Encyclopedia, 23, 865.

Duran, M. A., & Grossmann, I. E. (1986). Simultaneous optimizationand heat integration of chemical processes. AIChE Journal, 32,123.

Feinberg, M., & Hildebrandt, D. (1997). Optimal reactor design froma geometric viewpoint. Chemical Engineering Science, 52(10),1637.

Fjeld, M., Asbjornsen, O. A., & Astrom, K. J. (1974). Reactioninvariants and their importance in the analysis of eigenvectors,stability and controllability of CSTRs. Chemical Engineering Sci-ence, 30, 1917.

Gel’Perin, E. I., Bakshi, Y. M., Avetisov, A. K., & Gel’bshtein, A. I.(1979). Kinetic model of the oxidative chlorination of ethane tovinyl chloride II. Kinetica I Kataliz, 20, 129.

Gel’Perin, E. I., Bakshi, Y. M., Avetisov, A. K., & Gel’bshtein, A. I.(1983). Kinetic model of the oxidative chlorination of ethane tovinyl chloride III. Kinetica I Kataliz, 24, 633.

Gel’Perin, E. I., Bakshi, Y. M., Avetisov, A. K., & Gel’bshtein, A. I.(1984). Kinetic model of the oxidative chlorination of ethane tovinyl chloride IV. Kinetica I Kataliz, 25, 842.

Glasser, D., Crowe, C., & Hildebrandt, D. (1803). A geometricapproach to steady flow reactors: the attainable region and opti-mization in concentration space. Industrial Engineering ChemimalResearch, 26(9), 1987.

Hildebrandt, D., Glasser, D., & Crowe, C. (1990). The geometry ofthe attainable region generated by reaction and mixing with andwithout constraints. Industrial Engineering Chemical Research,29(1), 49.

Horn, F. J. M. (1964). Attainable regions in chemical reactiontechnique. In The Third European Symposium on Chemical Reac-tion Engineering. London: Pergamon.


Horn, F. J. M., & Tsai, M. J. (1987). The use of adjoint variables inthe development of improvement criteria for chemical reactors.Journal of Optimization Theory and Applications, 1(2), 131.

Jackson, R. (1968). Optimization of chemical reactors with respect toflow configuration. Journal of Optimization Theory and Applica-tions, 2(4), 240.

Karra, S. B., & Senkan, S. M. (1988). A detailed chemical kineticmechanism for the oxidative pyrolysis of CH3Cl. Industrial Engi-neering Chemical Research, 27, 1163.

Kokossis, A. C., & Floudas, C. A. (1990). Optimization of complexreactor networks–I. isothermal operation. Chemical EngineeringScience, 45(3), 595.

Kokossis, A.C., & Floudas, C.A. Synthesis of non-isothermal reactornetworks. Annual AIChE meeting, San Francisco, CA (1991).

Kurtz, B. E. (1972). Homogeneous kinetics of methyl chloride chlori-nation. Industrial Engineering Chemical Research, 11, 332.

Lakshmanan, A., & Biegler, L.T. (1995). Reactor network targetingfor waste minimization. In M. El-Halwagi and D. Petrides (eds)Pollution Pre6ention 6ia Process and Product Modifications,AIChE Symposium Series, Vol. 90, p. 128.

Lakshmanan, A., & Biegler, L. T. (1996a). Synthesis of optimalchemical reactor networks. Industrial Engineering Chemical Re-search, 35(4), 1344.

Lakshmanan, A., & Biegler, L. T. (1996b). Synthesis of optimalchemical reactor networks with simultaneous mass integration.Industrial Engineering Chemical Research, 35(12), 4523.

McPherson, R. W., Starks, C. M., & Fryar, G. J. (1979). Vinylchloride monomer… What you should know. Hydrocarbon Pro-cessing, 75, 75.

Omtveit, T., Tanskanen, J., & Lien, K. (1994a). Graphical targetingprocedures for reactor systems. Computers and Chemical Engi-neering, 18, S113.

Omtveit, T., Wahl, P. E., & Lien, K. M. (1994b). Decomposed

algorithmic synthesis of reactorseparation-recycle systems. Com-puters and Chemical Engineering, 18(11–12), 1115–1124.

Papalexandri, K. P., & Pistikopoulos, E. N. (1996). Generalizedmodular representation framework for process synthesis. AIChEJournal, 42(4), 1010.

Pibouleau, L. Floquet, & Domenech, S. (1988). Optimal synthesis ofreactor separator systems by nonlinear programming method.AIChE Journal, 34, 163.

Pro/II Casebook. Vinyl Chloride Monomer Plant. Simulation Sci-ences Inc. (1992).

Ranzi, E., Dente, M., Tiziano, F., Mullick, S., & Bussani, G. (1990).Mechanistic modeling of chlorinated reacting systems. Chimica el ’Industria (Milan), 72, 905.

Ranzi, E., Grottoli, M. G., Bussani, G., Che, S. C., & Zahng, G.(1993). A new simulation program predicts EDC furnace perfor-mances. Chimica e l ’Industria (Milan), 75, 261.

Smith, E. M. B., & Pantelides, C. C. (1995). Design of reaction/sepa-ration networks using detailed models. Computers Chemical Engi-neering, 19, S83–S88.

Turkay, M., & Grossmann, I.E. Structural flowsheet optimizationwith complex investment cost functions. paper 90a. Annual AIChEmeeting, Chicago, IL, 1996.

Viswanathan, J. V., & Grossmann, I. E. (1990). A combined penaltyfunction and outer approximation method for MINLP optimiza-tion. Computers and Chemical Engineering, 14, 769.

Wachi, S., & Morikawa, H. (1986). Liquid phase chlorination ofethylene and 1,2-dichloroethane. Journal of Chemical Engineeringof Japan, 19, 437.

Weissman, M., & Benson, S. W. (1984). Pyrolysis of methyl chloride.International Journal of Chemical Kinetics, 16, 307.

Yee, T. F., Grossmann, I. E., & Kravanja, Z. (1990). Simultaneousoptimization models for heat integration—II. heat exchangernetwork synthesis. Computers in Chemical Engineering, 14, 1165.

..

Reactor Network Synthesis Vinyl Chloride Proces

Documents

Transcript of Reactor Network Synthesis Vinyl Chloride Proces