towards further internal heat integration in design of reactive distillation

Chemical Engineering Science 61 (2006) 5377–5392www.elsevier.com/locate/ces

Towards further internal heat integration in design of reactive distillationcolumns—Part II. The process dynamics and operation

Kejin Huanga, Masaru Nakaiwaa,∗, Atsushi Tsutsumib

aResearch Institute for Innovation in Sustainable Chemistry, National Institute of Advanced Industrial Science and Technology (AIST), Tsukuba 305-8565, JapanbDepartment of Chemical System Engineering, The University of Tokyo, Tokyo 113-8565, Japan

Received 20 September 2004; received in revised form 30 November 2005; accepted 10 March 2006Available online 17 March 2006

Abstract

In the first paper of this series, it has been demonstrated that the capital investment and operating cost can frequently be reduced substantiallythrough seeking further internal heat integration between the reaction operation and separation operation for a reactive distillation columninvolving reactions with highly thermal effect. In this paper, the dynamics and operation of the resultant reactive distillation system is to beexamined, with special emphasis focused on the dynamic effect of the supplementary internal heat integration. It has been found that seekingfurther internal heat integration can sometimes improve process dynamics and lessen difficulties in process operation. This outcome stemsfrom the refined relationship between the reaction operation and separation operation involved and is of great significance in tightening processdesign for a reactive distillation column containing reactions with highly thermal effect.

It should, however, be pointed out that seeking further internal heat integration might also confine severely the flexibility of the resultantreactive distillation column due to the reduction of mass transfer driving forces. When encountering a sharp increase in the product specificationthat is more relevant to the supplementary internal heat integration, the process might show deteriorated dynamic performance and can evenconverge to an undesirable steady state where the economical advantages of the supplementary internal heat integration are lost totally. Therefore,some effective measures to increase the redundancy of the resultant process design have to be taken to deal with the side-effect during processdevelopment.� 2006 Elsevier Ltd. All rights reserved.

Keywords: Distillation; Reaction engineering; Heat integration; System engineering; Process control; Dynamic simulation

1. Introduction

For a reactive distillation column containing reactions withhighly thermal effect, it has already been demonstrated that asubstantial reduction of operating cost and capital investmentcan frequently be achieved by seeking further internal heat in-tegration between the reaction operation and separation opera-tion involved (Huang et al., 2005). Although the improvementin process design looks very attractive, the influences to processdynamics and operation remain a matter of special concern. Asthe combination of the reaction operation and separation oper-ation already causes a reactive distillation column more diffi-cult to control than a conventional distillation column, it will

∗ Corresponding author. Tel.: +81 029 861 4696; fax: +81 029 861 4660.E-mail address: [email protected] (M. Nakaiwa).

0009-2509/$ - see front matter � 2006 Elsevier Ltd. All rights reserved.doi:10.1016/j.ces.2006.03.015

certainly perplex process designers if very complicated processdynamics and severe operation difficulties can arise from thisprocess innovation. It seems therefore imperative to investigatethe detailed impact from the reinforcement of internal heat in-tegration between the reaction operation and separation opera-tion on process dynamics and operation.

Thanks to the tremendous efforts exerted on reactive dis-tillation systems worldwide so far, a deep insight has alreadybeen acquired into the process dynamics and operation. A greatnumber of papers addressed the descriptions of reactive distil-lation operation by various approaches, which could generallybe categorized into equilibrium and non-equilibrium cell mod-els (Roat et al., 1986; Cuille and Reklaitis, 1986; Ruiz et al.,1995; Alejski and Duprat, 1996; Kreul et al., 1998; Baur et al.,2001; Peng et al., 2003). The possible existence of multiplesteady states sparked a considerable interest in the mecha-nism behind this intricate phenomenon (Jacobs and Krishna,

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5378 K. Huang et al. / Chemical Engineering Science 61 (2006) 5377–5392

1993; Ciric and Miao, 1994; Hauan et al., 1995; Sneesby et al.,1998; Eldarsi and Douglas, 1998; Mohl et al., 1999; Guttingerand Morari, 1999a,b). The uncertainties introduced by the inputand/or output multiplicities provoked serious concern on thesmooth and reliable process operation and thus stimulated thedevelopment of robust control systems to deal with such com-plexities (Kumar and Daoutidis, 1999; Rosendo et al., 2000;Vora and Daoutidis, 2001). For the tight design of control sys-tems, the complex process dynamics and strong nonlinearityhad to be well compensated and this generally necessitated del-icate modeling techniques and complicated control algorithms(Bisowarno et al., 2004). Recently, nonlinear model-based con-trol algorithms were also employed to the operation of somereactive distillation columns (Balasubramhanya and Doyle III,2000; Tian et al., 2003; Engell and Fernholz, 2003). Gener-ally speaking, for the effective operation of a reactive distilla-tion column, maintaining tightly product qualities and keepinga strict stoichiometric balance between reactants are the twoessential objectives in the control system design. Although theproduct qualities can be controlled, in most cases, in a wayexactly similar to the control of a conventional distillation col-umn (e.g., by temperature or direct composition control), theproper control of the reactant stoichiometry appears to be quitetroublesome. Ratio control system appears to be an appropriatecandidate, however, the poor accuracy of flow measurementsprohibits it from being an effective one (Chiang et al., 2002).Sneesby et al. (1999) proposed a two-point control scheme foran ethyl tert-butyl ether (ETBE) reactive distillation column inwhich both the product purity and the conversion rate couldbe controlled. An inferential model had, however, to be devel-oped and employed to detect the degree of conversion withinthe reactive section. Luyben and his coworkers (Luyben, 2000;Al-Arfaj and Luyben, 2000) proposed and evaluated six alter-native control structures for a hypothetical reactive distillationcolumn. The use of a concentration analyzer in the reactive sec-tion was advocated to keep the stoichiometric balance betweenreactants. They further proved the necessity and the effective-ness of the internal composition control loop in the operationof an ETBE and a tert-amyl methyl ether (TAME) reactivedistillation columns (Al-Arfaj and Luyben, 2002, 2004). Wanget al. (2003a,b) proposed a feed ratio plus internal concentrationcontrol scheme and found effective for two reactive distillationcolumns producing methyl tertiary butyl ether (MTBE) andn-butyl acetate, respectively. Except for the feedforward com-pensation effect involved, their control configuration workedactually in the same principle as that of Al-Arfaj and Luyben.Huang et al. (2004) devised a temperature plus feed ratio cas-cade control scheme for two heterogeneous reactive distillationcolumns synthesizing n-butyl propionate and butyl acetate, re-spectively. In order to deal with the variations in the productionrate, they adjusted the setpoint of the temperature controller ac-cording to a simple regression model developed. With regardto process design and its influences to process dynamics andoperation, only a few papers have appeared on reactive distil-lation systems so far. Heath et al. (2000) discussed the selec-tion of control structures in terms of an economical objectivefunction for an ethylene glycol reactive distillation column.

Georgiadis et al. (2002) addressed the interaction between pro-cess design and control for a reactive distillation column syn-thesizing ethyl acetate and advocated the simultaneous consid-eration of process design and process operation in the earlystage of process development. A common feature of these stud-ies is the employment of a mixed-integer nonlinear program-ming (MINLP) approach. Although it is a very powerful syn-thesis method, it does not offer a clear picture that can reflectthe intricate relationship between process design and processdynamics and operation. For design of an intensified chemi-cal process like a reactive distillation column, it is very helpfulfor a designer to know clearly some favorable design optionsand their detailed effect on process dynamics and operation inorder to make an economical and yet reliable decision. There-fore, how to effectively tighten process design and meanwhileleave as small a negative effect as possible and even a positiveeffect to process dynamics and operation still remains to be animportant issue to solve in the near future.

The purpose of the present work is to evaluate the dynamiceffect of the supplementary internal heat integration betweenthe reaction operation and separation operation for a reactivedistillation column involving reactions with highly thermal ef-fect. After a brief introduction of the principle of internal heatintegration, the process designs for two reactive distillationcolumns containing, respectively, a highly exothermic reactionand a highly endothermic one are described. Studies of processdynamics and operation are then performed on these two pro-cesses. Comparison of open-loop and closed-loop dynamic be-haviors is conducted in depth between the process designs withand without further internal heat integration between the re-action operation and separation operation. Transitions to someundesirable steady states have been encountered in the servoresponses and found to hold a close relationship with the sup-plementary internal heat integration. The implications of the re-inforcement of internal heat integration are discussed and someconcluding remarks are highlighted in the last section of thepaper.

2. Process design by seeking further internal heatintegration

For a reactive distillation column containing reactions withhighly thermal effect, the appropriate superimposition of re-active section onto rectifying/stripping section can serve tostrengthen internal heat integration between the reaction op-eration and separation operation involved. There usually existtwo design options in the process development. One is to ex-tend reactive section to either rectifying section (in case of en-dothermic reactions) or stripping section (in case of exothermicreactions) in terms of the detailed reaction system at hand. Theother is to choose deliberately the feed location of the relevantreactant in the reactive section. Starting from a basic processdesign with three distinct sections (i.e., rectifying, reactive andstripping sections), one can easily evolve it into a more effi-cient one with the combinatorial utilization of these two meth-ods. More detailed information can be found in the first paperof this series.

K. Huang et al. / Chemical Engineering Science 61 (2006) 5377–5392 5379

Table 1Physical properties and nominal operating condition of hypothetical reactivedistillation columns

Items Systems

Example I Example II

Pressure (Bar) 9 9Overhead product compo-sition (C, mole fraction)

0.95 0.95

Bottom product composi-tion (D, mole fraction)

0.95 0.95

Holdup (kmol) Stage 1 1Condenser 20 20Reboiler 20 20

Activation energy(kJ kmol−1)

Forward 125 520 167 360

Backward 167 360 125 520

Specific reaction rate at366 K (kmol s−1 kmol−1)

Forward 0.008 0.008

Backward 0.004 0.004Relative volatilityA:B:C:D

4:2:8:1 4:2:8:1

Heat of reaction(kJ kmol−1)

−41 840 +41 840

Latent heat of vaporiza-tion (kJ kmol−1)

29053.7 29053.7

Vapor pressureconstants

A (Avp/Bvp) 12.3463/3862 12.3463/3862

B (Avp/Bvp) 11.6531/3862 11.6531/3862C(Avp/Bvp) 13.0394/3862 13.0394/3862D(Avp/Bvp) 10.96/3862 10.96/3862

The two hypothetical reactive distillation systems, i.e.,examples I and II, studied in the first paper of this series,are employed here for examining the dynamic effect of thesupplementary internal heat integration between the reactionoperation and separation operation. In the reactive section ofthese reactive distillation columns, reactants A and B undergoa reversible liquid-phase reaction to form products C and D

A + B ↔ C + D. (1)

The volatilities are such that the products C and D are the light-est and heaviest, respectively, in the system. The net reactionrate for component i on stage j in the reactive section is given by

ri,j = viHj (Kf,j xA,j xB,j − Kb,j xC,j xD,j ), (2.1)

where Kf,j and Kb,j are the forward and backward specificreaction rates and given by

Kf,j = �f e−Ef /(RT j ), (2.2)

Kb,j = �be−Eb/(RT j ). (2.3)

Ideal vapor and liquid phase behavior is assumed for the reac-tion system and the vapor–liquid equilibrium relationship canbe expressed as

Pj = xA,jPsA + xB,jP

sB + xC,jP

sC + xD,jP

sD, (3.1)

d = 0.0126 kmol/sxC, d = 0.95

FA = 0.0126 kmol/s

FB = 0.0126 kmol/s

RR = 0.033 kmol/s

Q = 822.86 kW14

9

2

21

b = 0.0126 kmol/sxD, bot = 0.95

(a)

(b)

d = 0.0126 kmol/sxC, d = 0.95

FB = 0.0126 kmol/s

FA = 0.0126 kmol/s

RR = 0.0286 kmol/s

Q = 695.942 kW17

13

9

2

21

b = 0.0126 kmol/sxD, bot = 0.95

14

Fig. 1. Process designs with and without further internal heat integration forexample I: (a) 7/6/7; (b) 7/6(1)/7(3).

yi,j = xi,jPsi /Pj . (3.2)

The vapor saturation pressure is calculated as

Ln P si = Avp,i − Bvp,i/Tj . (4)

Example I differs mainly from example II in the reaction ki-netics. The former contains a highly exothermic reaction andthe latter a highly endothermic one, both with an equal reactionheat. The detailed physical properties and nominal operatingcondition are shown in Table 1. Figs. 1 and 2 illustrate the ba-sic process designs, 7/6/7s, and those with further internal heatintegration between the reaction operation and separation oper-ation, 7/6(1)/7(3) and 7(3)/6(2)/7, for examples I and II, respec-tively. Here, a unified notation, Nr(n1)/Nrea(n2)/Ns(n3), hasbeen used to represent different process designs with and with-out further internal heat integration. Nr , Nrea, and Ns signify,respectively, the number of stages in the rectifying, reactive andstripping sections of the basic process design. The numbers inthe parentheses, n1 and n3, stand for the superimposition ofadditional reactive stages onto the rectifying/stripping section,respectively, and n2, denotes the movement of the feed location


d = 0.0126 kmol/sxC, d = 0.95

FA = 0.0126 kmol/s

FB = 0.0126 kmol/s

RR = 0.0142 kmol/s

Q = 1278.026 kW14

9

2

21

b = 0.0126 kmol/sxD, bot = 0.95

(a)

d = 0.0126 kmol/sxC, d = 0.95 FB = 0.0126 kmol/s

FA = 0.0126 kmol/s

RR = 0.0095 kmol/s

Q = 1142.787 kW

11

9

14

6

2

21

b = 0.0126 kmol/sxD, bot = 0.95

(b)

Fig. 2. Process designs with and without further internal heat integration forexample II: (a) 7/6/7; (b) 7(3)/6(2)/7.

of the relevant reactant within the reactive section. It is readilyseen that seeking further internal heat integration has lead to asimultaneous reduction of capital investment and energy con-sumption. In the remainder of the paper, the dynamics and op-eration are to be examined for the reactive distillation columnswith and without further internal heat integration between thereaction operation and separation operation. It is stipulated herethat the solid lines represent the dynamic responses of the ba-sic process designs, 7/6/7s, and the dashed lines the dynamicresponses of the process designs, 7/6(1)/7(3) and 7(3)/6(2)/7.

3. Example I: a reactive distillation system involving ahighly exothermic reaction

3.1. Open-loop transient responses to step changes in thefeed flow rate

In Fig. 3, the open-loop transient responses of the reactivedistillation columns with and without further internal heat in-

0.94

0.95

0.96

0.97

0.98

0.99

1

1.01

x C, d

(m

.f.)

-10%

+10%

0

0.2

0.4

0.6

0.8

1

1.2

0 2 4 6 8 10

Time [h]

0 2 4 6 8 10

Time [h]

x D,

bot (

m.f

.)

-10%

+10%

Fig. 3. Open-loop transient responses of the reactive distillation columns withand without further internal heat integration when they are subject to a ±10%step change in the feed flow rate of reactant B, respectively (example I).

tegration between the reaction operation and separation opera-tion are depicted, when they are subject to a ±10% step changein the feed flow rate of reactant B, respectively. After dis-turbed by a −10% step change in the feed flow rate of reactantB, the basic process design, 7/6/7, operates around the nomi-nal steady state for only about an hour and then drifts away.Notice that the new steady state reached is far different fromthe nominal one with a steady state composition nearly purefor component C in the top product and close to zero for com-ponent D in the bottom product, implying, therefore, an ab-normal situation with an extremely low conversion rate in thereactive section. Because of the slow composition dynamics,the sudden reduction in the feed flow rate of reactant B intro-duces no sharp changes in the composition profile within thereactive section and this is why the reactive distillation columncan exhibit relatively small deviations from the nominal steadystate for about an hour. The lack of reactant B leads to decre-ment in the conversion rate and affects gradually the internalvapor and liquid flow rates, thereby forcing the process to apartfrom the nominal operating condition. The variation of internalvapor and liquid flow rates accelerates further the runawayspeed, as is clearly indicated by the turning point around thetime of 3.2 h. The top product flow rate finally approaches tozero and makes the inventory control system of the reflux-drumcease to function properly. For the process design with furtherinternal heat integration between the reaction operation andseparation operation, 7/6(1)/7(3), it is readily seen that the sen-sitivity to the feed flow rate disturbance has been suppressed,substantially. In addition to a much slower escaping speed, the


0.94

0.945

0.95

0.955

0.96

0.965

0.97

0.975

0.98

0 2 4 6 8 10

Time [h]

0 2 4 6 8 10

Time [h]

x C, d

(m

.f.)

x D, b

ot (

m.f

.)

+0.0001 kmol/s

-0.0001 kmol/s

0.84

0.86

0.88

0.9

0.92

0.94

0.96

0.98

+0.0001 kmol/s

-0.0001 kmol/s

Fig. 4. Open-loop transient responses of the reactive distillation columnswith and without further internal heat integration when they are subject to a±0.0001 kmol/s step change in the reflux flow rate, respectively (example I).

steady state composition of component D remains around 0.26in the bottom product and is actually well above the corre-sponding value of the basic process design, 7/6/7, indicating arelatively higher conversion rate in the reactive section. Thereis no doubt that the improvement in process dynamics stemsfrom the reinforcement of internal heat integration between thereaction operation and separation operation. It will certainly befavorable to process operation.

When a +10% step change in the feed flow rate of reactantB has been encountered, fairly small differences have beenobserved between the process designs with and without furtherinternal heat integration between the reaction operation andseparation operation. Similar observations have also been madeon the open-loop transient responses to the step changes in thefeed flow rate of reactant A.

3.2. Open-loop transient responses to step changes in thereflux flow rate

The open-loop transient responses of the reactive distillationcolumns with and without further internal heat integration be-tween the reaction operation and separation operation are plot-ted in Fig. 4, after they are subject to a ±0.0001 kmol/s stepchange in the reflux flow rate, respectively. For the negative per-turbation in the reflux flow rate, there appears almost no sharpdifference between the process designs, 7/6/7 and 7/6(1)/7(3),whereas for the positive perturbation in the reflux flow rate,the former displays a much larger deviation than the latter inthe top and the bottom products, thereby implying a significant

0.94

0.95

0.96

0.97

0.98

0 42 6 8 10

Time [h]

0 42 6 8 10

Time [h]

x C, d

(m

.f.)

+0.0001 k mol/s

-0.0001 k mol/s

0.82

0.84

0.86

0.88

0.9

0.92

0.94

0.96

0.98

x D,

bot (

m.f

.)

-0.0001 k mol/s

+0.0001 k mol/s

Fig. 5. Open-loop transient responses of the reactive distillation columnswith and without further internal heat integration when they are subject to a±0.0001 kmol/s step change in the boilup flow rate, respectively (example I).

departure from the nominal steady state. Although the refluxflow rate has increased by only a small magnitude, it triggers aconsiderable change of the conversion rate in the reactive sec-tion, which is actually responsible for the runaway from thenominal operating condition. As the reflux flow rate is the ma-nipulative variable for the control of the top product, the greatdegree of asymmetry between the positive and the negative re-sponses (i.e., process nonlinearity) may present severely detri-mental influences to the control system performance. With theintroduction of further internal heat integration between the re-action operation and separation operation in the process design,7/6(1)/7(3), the degree of the process asymmetry has been al-leviated considerably. As can be seen in Fig. 4, the process isnow able to operate in the vicinity of the nominal steady stateand exhibits almost no noticeable differences between the pos-itive and the negative responses, therefore signifying a sharpimprovement in process dynamics.

3.3. Open-loop transient responses to step changes in thereboiler duty

Fig. 5 shows the open-loop transient responses of the reac-tive distillation columns with and without further internal heatintegration between the reaction operation and separation oper-ation, when they are subject to a ±0.0001 kmol/s step changein the boilup flow rate, respectively. No sharp difference hasbeen found between the process designs, 7/6/7 and 7/6(1)/7(3),in the presence of the positive perturbation in the boilup flow


rate. However, a large degree of difference has been observedin the presence of the negative perturbation in the boilup flowrate. The basic process design, 7/6/7, exhibits again a very largeshift from the nominal operating condition in a way extremelysimilar to the situation that a positive perturbation in the refluxflow rate is confronted. With the consideration of the supple-mentary internal heat integration between the reaction opera-tion and separation operation, the process design, 7/6(1)/7(3),is now able to work around the nominal steady state and thedegree of the asymmetrical nonlinearity has been relieved sub-stantially in comparison with that of 7/6/7. Since the boilupflow rate is the manipulative variable for the control of thebottom product, it is not difficult to understand that seekingfurther internal heat integration will be beneficial to processoperation.

3.4. Controllability analysis

Here, controllability analysis is conducted in the frequencydomain in terms of two predictive indexes: Morari resiliencyindex (MRI) and condition number (CN). The MRI is the min-imum singular value of the open-loop transfer function, whichcorresponds actually to a specific input and output direction.The control system that presents a large MRI over the fre-quency range of interest is preferred. The CN is the ratio ofthe maximum singular value to the minimum singular value,which can work as an effective measure for system sensitivityunder uncertainties in process parameters and modeling errors.The control system with a small CN over the frequency range

0

5

10

15

20

25

0.01 0.1 1 10 100 1000

ω [rad/s]

ω [rad/s]

MR

I

0

300

600

900

0.01 0.1 1 10 100 1000

CN

Fig. 6. Positive effect of the supplementary internal heat integration on processcontrollability (example I).

FC LC

FB

d, xC, d

LC

b, xD, bot CC

CC

FA

CC

9

2

21

14

Fig. 7. Control scheme for the reactive distillation column.

of interest is preferred. These two controllability indexes werefrequently applied to the evaluations of control system perfor-mance, especially in the development of thermodynamicallyefficient chemical processes through heat and/or mass integra-tion (Shimizu et al., 1985; Skogestad et al., 1990; Huang et al.,2000; Jimenez et al., 2001; Serra et al., 2003).

In Fig. 6, the dynamic effect of seeking further internal heatintegration between the reaction operation and separation oper-ation is demonstrated in terms of MRI and CN in the frequencydomain. As can be seen, after the reinforcement of internal heatintegration between the reaction operation and separation op-eration, the MRI has increased by a factor of 20 in the lowfrequency range. Even in the high frequency range, it has in-creased approximately by a factor of 3, although this tendencymay not be identified clearly from the illustration. As far as theCN is concerned, a substantial reduction has resulted from theintensification of internal heat integration between the reactionoperation and separation operation. It has contracted to about10 in the process design, 7/6(1)/7(3), from the correspondingvalues of the basic process design, 7/6/7, i.e., around 800 inthe low frequency range and 140 in the high frequency range.These outcomes predict obviously that seeking further internalheat integration is likely to present positive influences to pro-cess dynamics and thus eases process operation for the reactivedistillation system studied.

3.5. Closed-loop simulation

Closed-loop simulation is conducted to further ascertainthe dynamic effect of the supplementary internal heat integra-tion between the reaction operation and separation operation.Al-Arfaj and Luyben (2000) once devised six control structuresfor the reactive distillation column, and we select the controlscheme termed as CS1 by them in the following study. Thecontrol structure is reproduced in Fig. 7, where the purities ofboth the top and the bottom products are measured and con-trolled. In the top product, the composition of component C iscontrolled by manipulating the reflux flow rate. In the bottomproduct, the composition of component D is controlled by


0.93

0.94

0.95

0.96

0.97

0 1 2 3 4 5 6

Time [h]

0 1 2 3 4 5 6

Time [h]

0 1 2 3 4 5 6

Time [h]

0 1 2 3 4 5 6

Time [h]

x C, d

(m

.f.)

0.024

0.028

0.032

0.036

0.04

RR

[km

ol/s

]

0.94

0.944

0.948

0.952

0.956

x D, b

ot (

m.f

.)

0.02

0.022

0.024

0.026

0.028

0.03

0.032

Vnt

[km

ol/s

]

Fig. 8. Servo responses of the reactive distillation columns with and without further internal heat integration when the top control loop is subject to a ±0.01step change in setpoint, respectively (example I). Grey curves: 0.95 → 0.94; black curves: 0.95 → 0.96.

0.925

0.9375

0.95

0.9625

0.975

0 41 52 3 6

Time [h]

0 41 52 3 6

Time [h]

0 41 52 3 6

Time [h]

0 41 52 3 6

Time [h]

x C, d

(m

.f.)

+20%

-20%

0.015

0.025

0.035

0.045

RR

[km

ol/s

]

+20%

-20%

0.935

0.94

0.945

0.95

0.955

0.96

0.965

x D, b

ot (

m.f

.)

+20%

-20%

0.015

0.02

0.025

0.03

0.035

0.04

Vnt

[km

ol/s

] +20%

-20%

Fig. 9. Regulatory responses of the reactive distillation columns with and without further internal heat integration when production rate is subject to a ±20%step change, respectively (example I).


0.94

0.945

0.95

0.955

0.96

0.965

0 2 41 3 5 6

Time [h]

0 2 41 3 5 6

Time [h]

0 2 41 3 5 6

Time [h]

0 2 41 3 5 6

Time [h]

x C, d

(m

.f.)

0.02

0.03

0.04

RR

[km

ol/s

]

0.949

0.951

0.953

0.955

0.957

x D, b

ot (

m.f

.)

0.018

0.022

0.026

0.03

Vnt

[km

ol/s

]

Fig. 10. Regulatory responses of the reactive distillation columns with and without further internal heat integration when FB is changed from a pure componentflow of reactant B (zA,9 = 0.0 and zB,9 = 1.0) into a mixture flow of reactants A and B (zA,9 = 0.2 and zB,9 = 0.8) (example I).

manipulating the heat duty of reboiler. The concentration of re-actant A on stage 14 in the reactive section is measured and con-trolled by manipulating the feed flow rate of reactant A (i.e., theso-called reactant stoichiometry control loop). The feed flowrate of reactant B is the production rate handle and is flow con-trolled. Because the production rate is strictly controlled in thiscontrol scheme, a fair comparison is allowed between the pro-cess designs with and without further internal heat integrationbetween the reaction operation and separation operation. Thelevels of the reflux-drum and the bottom reboiler are controlledby the distillate and the bottom product flow rates, respectively.The dynamics of concentration measurements is assumed to betwo first-order lags of 30 s in series. The transmitter span of allcomposition measurements is taken to be 0.1 and all controlvalves are designed to be half open at the nominal steady state.

Proportional-only controllers are used for all level controlloops and proportional plus integral (PI) controllers are adoptedfor the top and the bottom composition control loops. Althoughit is reasonable to use a PI controller in the reactant stoichiome-try control loop, it may intensify the interaction between differ-ent control loops and thus worsen overall system performance.Since both the top and the bottom products are strictly con-trolled, the existence of a certain degree of discrepancy in thecomposition on stage 14 will not present strong influences tothe overall system performance. Hence, it seems more appro-priate to use a proportional-only controller here. All composi-tion controllers are tuned in terms of an experimental designapproach proposed by Finco et al. (1989). The ultimate gain

and the ultimate frequency are found by increasing the gain ofa proportional-only controller until sustained oscillations oc-cur. Then, the Ziegler–Nichols settings are calculated for eachcontrol loop. Finally, a detuning factor f is searched for thatmakes all composition control loops have appropriate dampingcoefficients

Kc = Kzn/f , (5)

TI = Tzn ∗ f , (6)

where Kc and TI represent the proportional gain and in-tegral time, respectively, and KZN and TZN denotes theZiegler–Nichols settings.

In Fig. 8, the dynamic responses of the reactive distillationcolumns with and without further internal heat integrationbetween the reaction operation and separation operation areillustrated, when the top control loop is subject to a ±0.01 stepchange in setpoint, respectively. It is readily to see that the pro-cess design with further internal heat integration, 7/6(1)/7(3),presents improved responses in comparison with the basic pro-cess design, 7/6/7, especially in the case when the setpoint isincreased from 0.95 to 0.96. In addition to the smoother tran-sitions between different specifications on the top product, theinteraction to the bottom control loop has also been relieved,confirming definitely the positive effect of the supplementaryinternal heat integration on process dynamics and operation.At the new steady states, the former can still maintain higherthermodynamic efficiency than the latter, thereby evidencing


the robustness of the supplementary internal heat integrationto the variations in operating conditions.

Fig. 9 shows the dynamic responses of the reactive distil-lation columns with and without further internal heat integra-tion between the reaction operation and separation operation,when the production rate is subject to a ±20% step change,respectively, in terms of a corresponding variation in the feedflow rate of reactant B. Again, it can be seen that the processdesign, 7/6(1)/7(3), outperforms 7/6/7, with relatively smallermaximum deviations and shorter setting time in both the topand the bottom control loops. The advantage in thermodynamicefficiency has been kept at the new steady states in spite of aconsiderable change in the production rate.

In Fig. 10, the dynamic responses of the reactive distillationcolumns with and without further internal heat integration be-tween the reaction operation and separation operation are dis-played, when FB is changed from a pure component flow ofreactant B (zA,9 = 0.0 and zB,9 = 1.0) into a mixture flow ofreactants A and B (zA,9 = 0.2 and zB,9 = 0.8). It should bereminded here of the fact that this represents one of the mostsevere situations to check the static and dynamic performanceof the supplementary internal heat integration. It is noted thatvery strong interaction has been observed between the top andthe bottom control loops in the basic process design, 7/6/7.With the reinforcement of internal heat integration between thereaction operation and separation operation in the process de-sign, 7/6(1)/7(3), far improved dynamic responses have beenobserved. The interaction has now been attenuated dramaticallybetween the top and the bottom control loops besides the re-duced setting time and maximum deviations. Even though aconsiderable change has taken place in the operating condition,seeking further internal heat integration still remains effectivein reducing energy consumption when compared with the basicprocess design, 7/6/7.

4. Example II: a reactive distillation system involving ahighly endothermic reaction

4.1. Open-loop transient responses to step changes in thefeed flow rate

In Fig. 11, the open-loop transient responses of the reac-tive distillation columns with and without further internal heatintegration between the reaction operation and separation op-eration are displayed, when they are subject to a ±10% stepchange in the feed flow rate of reactant A, respectively. Fromthe responses of the bottom product, it is hard to identify anydifferences between the process designs, 7/6/7 and 7(3)/6(2)/7.From the responses of the top product, a very large deviationhas been observed in the former than in the latter, in the pres-ence of the negative perturbation in the feed flow rate of reac-tant A. The gradual decline of the conversion rate in the reactivesection initiates a substantial change of internal vapor and liq-uid flow rates, thereby resulting in finally a considerable shiftfrom the nominal operating condition in the process design,7/6/7. With the introduction of further internal heat integrationbetween the reactive section and rectifying section in the pro-

0.45

0.6

0.75

0.9

1.05

0 42 6 8 10

Time [h]

0 42 6 8 10

Time [h]

x C, d

(m

.f.)

+10%

-10%

0.8

0.85

0.9

0.95

1

1.05

x D,

bot (

m.f

.)

-10%

+10%

Fig. 11. Open-loop transient responses of the reactive distillation columnswith and without further internal heat integration when they are subjectto a ±10% step change in the feed flow rate of reactant A, respectively(example II).

cess design, 7(3)/6(2)/7, the decline of the conversion rate hasbeen, to a certain extent, suppressed and the steady state devi-ation has been reduced in the top product. The low sensitivityto the feed flow rate disturbance is certainly attributed to thereinforcement of internal heat integration and can exercise abeneficial effect on process operation.

Fig. 12 illustrates the open-loop transient responses of thereactive distillation columns with and without further internalheat integration between the reaction operation and separationoperation, when they are subject to a ±10% step change inthe feed flow rate of reactant B, respectively. As can be seen,the responses obtained are extremely analogous to those thatare excited by the step changes in the feed flow rate of reactantA. Process design, 7(3)/6(2)/7, still exhibits a smaller steadystate deviation than 7/6/7 in the top product, confirming againthat the improvement in process dynamics originates from thereinforcement of internal heat integration between the reactionoperation and separation operation.

4.2. Open-loop transient responses to step changes in thereflux flow rate and reboiler duty

As far as the open-loop transient responses to step changesin the reflux and boilup flow rates are concerned, no signifi-cant differences have been found between the reactive distilla-tion columns with and without further internal heat integrationbetween the reaction operation and separation operation.


0.5

0.6

0.7

0.8

0.9

1

0 2 4 6 8 10

Time [h]

0 2 4 6 8 10

Time [h]

x C, d

(m

.f.)

-10%

+10%

0.84

0.87

0.9

0.93

0.96

0.99

1.02

x D,

bot (

m.f

.)

-10%

+10%

Fig. 12. Open-loop transient responses of the reactive distillation columnswith and without further internal heat integration when they are subjectto a ±10% step change in the feed flow rate of reactant B, respectively(example II).

4.3. Controllability analysis

The dynamic effect of seeking further internal heat integra-tion between the reaction operation and separation operation isagain investigated in terms of MRI and CN in the frequencydomain and the detailed results are illustrated in Fig. 13. At firstglance, seeking further internal heat integration looks to give anegative impact to control system performance. However, onemay find that it is not the case by a closer examination. Al-though the MRI has decreased from 14.1 to 10.9 and the CNhas increased from 6.8 to 8.8 in the low frequency range af-ter seeking further internal heat integration, the magnitudes ofchanges are actually very small, implying probably trivial oreven no degradation at all in system controllability. In the highfrequency range, almost no differences have been observed be-tween the process designs with and without further internal heatintegration, 7/6/7 and 7(3)/6(2)/7. Because process controlla-bility is more closely related to the high frequency behavior,it is therefore more reasonable to believe that seeking furtherinternal heat integration gives almost no influences to the con-trollability of the resultant process.

Recall the fact that seeking further internal heat integra-tion favors the rejection of disturbances from the feed flowrate of reactants A and B, it is not difficult to figure outthat the dynamic performance of the resultant process islikely to be improved as compared with the basic processdesign, 7/6/7.

0

5

10

15

20

0.01 0.1 1 10 100 1000

ω [rad/s]

0.01 0.1 1 10 100 1000

ω [rad/s]

MR

I0

5

10

15

CN

Fig. 13. Almost no effect of the supplementary internal heat integration onprocess controllability (example II).

4.4. Closed-loop simulation

The control system configuration is the same as is shownin Fig. 7. The dynamic responses of the reactive distillationcolumns with and without further internal heat integration be-tween the reaction operation and separation operation are shownin Fig. 14, when the bottom control loop is subject to a ±0.01step change in setpoint, respectively. As can be seen, a rathersmoother transition has been accomplished by the process de-sign, 7(3)/6(2)/7 than by 7/6/7, between different specificationson the bottom product, with not only smaller overshoots butalso shorter setting time. Furthermore, the interaction to thetop control loop appears to be alleviated, substantially. At thenew steady states reached the former can maintain its higherthermodynamic efficiency than the latter.

Fig. 15 depicts the dynamic responses of the reactive distil-lation columns with and without further internal heat integra-tion between the reaction operation and separation operation,when the production rate is subject to a ±10% step change,respectively. Process design, 7(3)/6(2)/7, exhibits again muchimproved dynamic responses than 7/6/7, with not only smallerdeviations but also shorter setting time. The advantage in ther-modynamic efficiency brought about by the supplementary in-ternal heat integration has been retained at the new steady statesreached.

In Fig. 16, the dynamic responses of the reactive distillationcolumns with and without further internal heat integration be-tween the reaction operation and separation operation are dis-played, when FA is changed from a pure component flow of


0.936

0.94

0.944

0.948

0.952

0.956

0 2 41 3 5 6

Time [h]

0 2 41 3 5 6

Time [h]

0 2 41 3 5 6

Time [h]

0 2 41 3 5 6

Time [h]

x C, d

(m

.f.)

0.008

0.01

0.012

0.014

0.016

RR

[km

ol/s

]0.93

0.94

0.95

0.96

0.97

x D, b

ot (

m.f

.)

0.034

0.04

0.046

0.052

Vnt

[km

ol/s

]

Fig. 14. Servo responses of the reactive distillation columns with and without further internal heat integration when the bottom control loop is subject to a±0.01 step change in setpoint, respectively (example II). Grey curves: 0.95 → 0.94; black curves: 0.95 → 0.96.

0.935

0.945

0.955

0.965

0 2 51 3 4 6

Time [h]

0 2 51 3 4 6

Time [h]

0 2 51 3 4 6

Time [h]

0 2 51 3 4 6

Time [h]

x C, d

(m

.f.)

-10%

+10%

0.008

0.011

0.014

0.017

RR

[km

ol/s

]

-10%

+10%

+10%

-10%

0.94

0.944

0.948

0.952

0.956

0.96

x D, b

ot (m

.f.)

+10%

-10%

0.03

0.035

0.04

0.045

0.05

Vnt

[km

ol/s

]

-10%

+10%

Fig. 15. Regulatory responses of the reactive distillation columns with and without further internal heat integration when production rate is subject to a ±10%step change, respectively (example II).

reactant A (zA,14 = 1.0 and zB,14 = 0.0) into a mixture flow ofreactants A and B (zA,14 = 0.95 and zB,14 = 0.05). Again, farstronger interaction has been observed between the top and the

bottom control loops in the basic process design, 7/6/7, thanin the process design, 7(3)/6(2)/7. Seeking further internal heatintegration can still secure a substantial reduction in energy


0.944

0.946

0.948

0.95

0.952

0 1 42 3 65

Time [h]

0 1 42 3 65

Time [h]

0 1 42 3 65

Time [h]

0 1 42 3 65

Time [h]

x C, d

(m

.f.)

0.008

0.01

0.012

0.014

0.016

RR

[km

ol/s

]0.945

0.9471

0.9492

0.9513

x D, b

ot (

m.f

.)

0.036

0.039

0.042

0.045

0.048

Vnt

[km

ol/s

]

Fig. 16. Regulatory responses of the reactive distillation columns with and without further internal heat integration when FA is changed from a pure componentflow of reactant A (zA,14 = 1.0 and zB,14 = 0.0) into a mixture flow of reactants A and B (zA,14 = 0.95 and zB,14 = 0.05) (example II).

consumption despite a considerable shift from the nominaldesign condition.

5. Discussions

Al-Arfaj and Luyben (2000) examined systematically thecontrol system design for example I and indicated that augmentof liquid holdups (to say more exactly, the amount of catalystemployed) in the reactive section could improve the dynamicperformance of the reactive distillation column. In terms of thepresent study, we may further systematize their conclusion:even though without increasing the amount of catalyst, thedynamic performance could still be enhanced by distributingappropriately the reactive section within the process. Not onlythe open-loop responses but also the closed-loop ones justifythe possibility. The extension of their conclusion signifies thevital importance of process design to process dynamics andoperation. More specifically, the improvement in process dy-namics and operation can sometimes result from the elabora-tion of process design with almost no additional capital cost fora reactive distillation system involving reactions with highlythermal effect.

It is worthwhile to look into the fact that seeking furtherinternal heat integration can present positive influences to thedynamics and operation of a reactive distillation column in-volving reactions with highly thermal effect. There is no doubtthat the combination of the reaction operation and separationoperation in a single unit leads to a more challenging problemin process operation than a conventional distillation column.The essential difficulties arise actually from how to arrange this

combination in process design. It is one of the most importantdesign objectives that affect not only process economics butalso process controllability in all kinds of process intensifica-tions. Simply distributing the reactive section between the rec-tifying and stripping sections (e.g., in a basic process design)cannot always be the best design option, especially for somereactions with highly thermal effect. On the contrary, pursuingan optimum combination of the reaction operation and separa-tion operation can sometimes gain simultaneous improvementin both aspects, because a more refined relationship has beendeveloped between the reaction operation and separation opera-tion. Seeking further internal heat integration serves to enhancethe synergism between the reaction operation and separationoperation and it is the primary reason that process economicsand process controllability have been improved simultaneouslyfor the two reactive distillation columns studied in this work.

It must be pointed out here that seeking further internal heatintegration between the reaction operation and separation oper-ation might also give severe restrictions to the flexibility of theresultant reactive distillation column. Figs. 17 and 18 presenta detailed comparison of the servo responses to different mag-nitudes of step perturbations in the setpoints of control loopsfor examples I and II. Here, the step perturbations are made onthe end products that are more relevant to the supplementaryinternal heat integration, i.e., the bottom product for example Iand the top product for example II. It is noted that for the basicprocess designs, 7/6/7s, there appear no problems at all in thetransitions between different product specifications regardlessof the magnitudes of the perturbations. However, for the pro-cess designs with further internal heat integration, 7/6(1)/7(3)


0.9

0.92

0.94

0.96

0.98

0 1 2 43 5 6

Time [h]

0 1 2 43 5 6

Time [h]

0 1 2 43 5 6

Time [h]

0 1 2 43 5 6

Time [h]

x C, d

(m

.f.)

0.025

0.03

0.035

0.04

0.045

0.05

RR

[km

ol/s

]

0.94

0.95

0.96

0.97

0.98

x D, b

ot (

m.f

.)

0.015

0.02

0.025

0.03

0.035

0.04

0.045

Vnt

[km

ol/s

]

Fig. 17. Servo responses of example I when the bottom control loop is subject to a +0.01 and a +0.02 step changes in setpoint, respectively. Grey curves:0.95 → 0.96; black curves: 0.95 → 0.97.

0.944

0.948

0.952

0.956

0.96

0.964

0 1 2 3 4 5 6

Time [h]

0 1 2 3 4 5 6

Time [h]

0 1 2 3 4 5 6

Time [h]

0 1 2 3 4 5 6

Time [h]

x C, d

(m

.f.)

0.008

0.01

0.012

0.014

0.016

RR

[km

ol/s

]

0.944

0.946

0.948

0.95

0.952

x D, b

ot (

m.f

.)

0.0378

0.0408

0.0438

0.0468

Vnt

[km

ol/s

]

Fig. 18. Servo responses of example II when the top control loop is subject to a +0.005 and a +0.01 step change in setpoint, respectively. Grey curves:0.95 → 0.955; black curves: 0.95 → 0.96.


Table 2Multiple steady states and the performance of reactive distillation systems

Systems Items

Process configurations Reflux flow rate (mol/s) Reboiler duty (kW) Features of steady states

Example I: xC,d = 0.95, xD,bot = 0.97 7/6/7 34.7287 871.46 Nominal7/6(1)/7(3) 31.037 764.17 Desirable7/6(1)/7(3) 41.6447 1167.88 Undesirable

Example II: xC,d = 0.96, xD,bot = 0.95 7/6/7 14.9466 1305.82 Nominal7(3)/6(2)/7 No solution No solution Desirable7(3)/6(2)/7 28.5072 1699.88 Undesirable

0.953

0.954

0.955

0.956

0.957

0.01 0.012 0.014 0.016 0.018

RR [kmol/s]

x C, d

(m

.f.)

Fig. 19. Relationship between the reflux flow rate and the top product com-position when the bottom product is controlled at 0.95 for 7(3)/6(2)/7 ofexample II.

and 7(3)/6(2)/7, their dynamic responses appear to be quitesensitive to the magnitudes of the perturbations. For the rela-tive small perturbations, they really exhibit better dynamic re-sponses than the basic process designs, 7/6/7s, whereas for therelative large perturbations, they display much degraded ones.As can be seen, both the top and the bottom control loops be-have very sluggishly after initial strong actions and leave largediscrepancies for an extremely long period of time. The PIcontrollers increase slowly the reflux and boilup flow rates inorder to eliminate the discrepancies and drive finally these pro-cesses to two undesirable steady states. Table 2 gives a compre-hensive comparison between the steady states for examples Iand II. It can be seen that the process designs with further in-ternal heat integration lose their advantages in thermodynamicefficiency over the basic process designs at those undesirablesteady states. Although there does exist a desirable steady statefor the process design, 7/6(1)/7(3), the process fails to workaround it during dynamic operation. For the process design,7(3)/6(2)/7, even no desirable steady state can be found andthis is well demonstrated by the relationship between the refluxflow rate and the top product composition shown in Fig. 19. Itindicates clearly that the top product cannot be concentrated to0.96 with a smaller reflux rate than that of the basic process de-sign, 7/6/7. The failure to reach those desirable steady states isapparently related to the reinforcement of internal heat integra-tion between the reaction operation and separation operationand therefore some effective measures must be taken to dealwith this side-effect during process design. Remember the fact

that seeking further internal heat integration reduces actuallythe driving force of mass transfer in the rectifying/stripping sec-tion where internal heat integration has been arranged (Melleset al., 2000), it certainly becomes more difficult to reach ahigh degree of separation in process designs, 7/6(1)/7(3) and7(3)/6(2)/7 than in 7/6/7s. This is why the transitions to uneco-nomical steady states occur in the dynamic operation. A directmethod to solve this problem is therefore to make appropriatecompensation for the loss in the driving forces of mass trans-fer by adding some separating stages in the corresponding sec-tion. An alternative method is to reduce somehow the degree ofinternal heat integration between the reaction operation and sep-aration operation. Although no requirement is needed on cap-ital investment, the degradation in thermodynamic efficiencyis inevitable. Therefore, the first method is generally preferredto the second one in process development. In Figs. 17 and 18,the effect of adding three separating stages in process designs,7/6(1)/10(3) and 10(3)/6(2)/7, is also demonstrated (representedhere by the dotted lines). It is readily seen that the deficien-cies associated with the supplementary internal heat integrationhave been remedied, completely. Not only have the transitionsto those undesirable steady states been avoided, but also thesuperiorities over the basic process designs, 7/6/7s, have beenmaintained in both static and dynamic responses.

6. Conclusion

In addition to a substantial reduction of capital investmentand operating costs, seeking further internal heat integrationbetween the reaction operation and separation operation cansometimes improve process dynamics and alleviate difficultiesin process operation for a reactive distillation column involv-ing reactions with highly thermal effect. As long as processdesign has been carried out adequately, the supplementaryinternal heat integration between the reaction operation andseparation operation has been found to be relatively insensitiveto the changes in operating conditions from the viewpoint ofprocess dynamics and operation and this is in good accordancewith the conclusion obtained in the first paper of this series.The reason can solely be attributed to the refined relationshipbetween the reaction operation and separation operation result-ing from the reinforcement of internal heat integration. Thesefindings emphasize the utmost importance of tightening pro-cess design by means of strengthening the synergistic effect


within a reactive distillation column. Although these conclu-sions have been derived from the two hypothetical reactivedistillation systems studied in this work, they are consideredto be of great significance and can be effective guidelines forthe design and operation of some reactive distillation systemsinvolving reactions with highly thermal effect.

However, it must be borne in mind that seeking further in-ternal heat integration between the reaction operation and sep-aration operation might also confine severely the flexibility ofthe resultant reactive distillation column due to the reductionof mass transfer driving forces. When a large increase has beenencountered in the product specification that is more relevantto the supplementary internal heat integration, the resultantprocess design might show much deteriorated dynamic perfor-mance and can even approach to an undesirable steady statewhere the economical advantages of the supplementary inter-nal heat integration have been lost totally. Therefore, some ef-fective measures, e.g., adjustment of the degree of internal heatintegration and addition of a number of separating stages inthe rectifying/stripping section, must be taken to ensure the re-sultant process design with enough redundancy and flexibilityduring process development. Current work is now underway todevelop a systematic method that permits feasible process de-sign for a reactive distillation column involving reactions withhighly thermal effect.

Notation

A reactive componentAvp vapor pressure constant, Pab bottom withdrawal, kmol s−1

B reactive componentBvp vapor pressure constant, Pa KC product componentCC composition controllerCN condition numberd distillate flow rate, kmol s−1

D product componentE activation energy of a reaction, kJ kmol−1

f detuning parameterF feed flow rate of reactants, kmol s−1

FC flow rate controllerH holdup, kmolK specific reaction rate, kmol s−1 kmol−1

Kc proportional gainL liquid flow rate, kmol s−1

LC level controllerMRI Morari resilience indexn number of stagesN number of stagesP pressure, kPaPI proportional plus integral controllerQ heat duty, kWr net reaction rate, kmol s−1

R ideal gas law constant, kJ kmol−1 K−1

RR reflux flow rate, kmol s−1

T temperature, KTI integral time, sV vapor flow rate, kmol s−1

x liquid compositiony vapor compositionz feed composition

Greek letters

� pre-exponential factorv stoichiometric coefficients of a reaction� frequency, rad s−1

Subscripts

A component indexb backward reactionbot bottom productB component indexC component indexd distillateD component indexf forward reactioni component indexj stage indexnt total number of stagesr rectifying sectionrea reactive sections stripping sectionZN Ziegler–Nichols settings

Superscripts

s saturation

Acknowledgments

A dynamic model of the hypothetical reactive distillationcolumn involving a highly exothermic reaction is provided byProf. W. L. Luyben and Prof. M. A. Al-Arfaj and hereby is ac-knowledged. The authors are grateful to the financial supportfrom the Japan Science and Technology (JST) Corporation un-der the frame of Core Research and Evolutional Science andTechnology (CREST).

References

Al-Arfaj, M.A., Luyben, W.L., 2000. Comparison of alternative controlstructures for an ideal two-product reactive distillation column. IndustrialEngineering and Chemistry Research 39, 3298–3307.

Al-Arfaj, M.A., Luyben, W.L., 2002. Control study of ethyl tert-butyl etherreactive distillation. Industrial Engineering and Chemistry Research 41,3784–3796.

Al-Arfaj, M.A., Luyben, W.L., 2004. Plantwide control for TAME productionusing reactive distillation. A.I.Ch.E. Journal 50, 1462–1473.

Alejski, K., Duprat, F., 1996. Dynamic simulation of the multi-componentreactive distillation. Chemical Engineering Science 51, 4237–4252.

Balasubramhanya, L.S., Doyle III, F.J., 2000. Nonlinear model-based controlof a batch reactive distillation column. Journal of Process Control 10,209–218.

Baur, R., Taylor, R., Krishna, R., 2001. Dynamic behavior of reactivedistillation columns described by a non-equilibrium stage model. ChemicalEngineering Science 56, 2085–2102.


Bisowarno, B.H., Tian, Y.C., Tade, M.O., 2004. Adaptive control of an ETBEreactive distillation column. Journal of Chemical Engineering of Japan 37,210–216.

Chiang, S.F., Kuo, C.L., Yu, C.C., Wong, D.S.H., 2002. Design alternatives forthe amyl acetate process: coupled reactor/column and reactive distillation.Industrial Engineering and Chemistry Research 41, 3233–3246.

Ciric, A.R., Miao, P., 1994. Steady state multiplicities in an ethylene glycolreactive distillation column. Industrial Engineering and Chemistry Research33, 2738–2748.

Cuille, P.E., Reklaitis, G.V., 1986. Dynamic simulation of multi-componentbatch rectification with chemical reactions. Computers and ChemicalEngineering 10, 389–398.

Eldarsi, H.S., Douglas, P.L., 1998. Methyl tert-butyl ether catalytic distillationcolumn: multiple steady states. Transactions of the Institution of ChemicalEngineers Part A 76, 509–516.

Engell, S., Fernholz, G., 2003. Control of a reactive separation process.Chemical Engineering and Processing 42, 201–210.

Finco, M.V., Luyben, W.L., Polleck, R.E., 1989. Control of distillationcolumns with low relative volatilities. Industrial Engineering and ChemistryResearch 28, 75–83.

Georgiadis, M.C., Schenk, M., Pistikopoulos, E.N., Gani, R., 2002. Theinteraction of design, control and operability in reactive distillation systems.Computers and Chemical Engineering 26, 735–746.

Guttinger, T.E., Morari, M., 1999a. Predicting multiple steady statesin equilibrium reactive distillation. I. Analysis of nonhybrid systems.Industrial Engineering and Chemistry Research 38, 1633–1648.

Guttinger, T.E., Morari, M., 1999b. Predicting multiple steady states inequilibrium reactive distillation. II. Analysis of hybrid systems. IndustrialEngineering and Chemistry Research 38, 1649–1665.

Hauan, S., Hertzberg, T., Lien, K.M., 1995. Why methyl tert-butyl etherproduction by reactive distillation may yield multiple solutions. IndustrialEngineering and Chemistry Research 34, 987–991.

Heath, J.A., Kookos, I.K., Perkins, J.D., 2000. Process control structureselection based on economics. A.I.Ch.E. Journal 46, 1998–2016.

Huang, K., Qian, J., Nakaiwa, M., Takamatsu, T., 2000. Assessment of controlconfigurations for a general heat-integrated distillation column. ChineseJournal of Chemical Engineering 8, 339–346.

Huang, S.G., Kuo, C.L., Hung, S.H., Chen, Y.W., Yu, C.C., 2004. Temperaturecontrol of heterogeneous reactive distillation. A.I.Ch.E. Journal 50,2203–2216.

Huang, K., Iwakabe, K., Nakaiwa, M., Tsutsumi, A., 2005. Towards furtherinternal heat integration in design of reactive distillation columns—Part I:the design principle. Chemical Engineering Science 60, 4901–4914.

Jacobs, R., Krishna, R., 1993. Multiple solutions in reactive distillation formethyl tert-butyl ether synthesis. Industrial Engineering and ChemistryResearch 32, 1706–1709.

Jimenez, A., Hernandez, S., Montoy, F.A., Zavala-Garcia, M., 2001. Analysisof control properties of conventional and non-conventional distillationsequences. Industrial Engineering and Chemistry Research 40, 3757–3761.

Kreul, L.U., Gorak, A., Dittrich, C., Barton, P.I., 1998. Dynamic catalyticdistillation: advanced simulation and experimental validation. Computersand Chemical Engineering 22, S371–S378.

Kumar, A., Daoutidis, P., 1999. Modeling, analysis and control of ethyleneglycol reactive distillation column. A.I.Ch.E. Journal 45, 51–68.

Luyben, W.L., 2000. Economic and dynamic impact of the use of excessreactant in reactive distillation systems. Industrial Engineering andChemistry Research 39, 2935–2946.

Melles, S., Grievink, J., Schrans, S.M., 2000. Optimization of the conceptualdesign of reactive distillation columns. Chemical Engineering Science 55,2089–2097.

Mohl, K.D., Kienle, A., Gilles, E.D., Rapmund, P., Sundmacher, K.,Hoffmann, U., 1999. Steady state multiplicities in reactive distillationcolumns for the production of fuel ethers MTBE and TAME: theoreticalanalysis and experimental verification. Chemical Engineering Science 54,1029–1043.

Peng, J., Edgar, T.F., Eldridge, R.B., 2003. Dynamic rate-based andequilibrium models for a packed reactive distillation column. ChemicalEngineering Science 58, 2671–2680.

Roat, S.D., Downs, J.J., Vogel, E.F., Doss, J.E., 1986. The integration ofrigorous dynamic modeling and control system synthesis for distillationcolumns: an industrial approach. In: Morari, M., McAvoy, T.J. (Eds.),Chemical Process Control—CPC III. Elsevier, New York.

Rosendo, M.L., Eduardo, P.C., Jose, A.R., 2000. A robust PI controlconfiguration for a high-purity ethylene glycol reactive distillation column.Chemical Engineering Science 55, 4925–4937.

Ruiz, C.A., Basualdo, M.S., Scenna, N.J., 1995. Reactive distillation dynamicsimulation. Transactions of the Institution of Chemical Engineers Part A73, 363–378.

Serra, M., Espuna, A., Puigjaner, L., 2003. Controllability of different multi-component distillation arrangements. Industrial Engineering and ChemistryResearch 42, 1773–1782.

Shimizu, K., Holt, B.R., Morari, M., Mah, R.S.H., 1985. Assessment ofcontrol structures for binary distillation columns with secondary refluxand vaporization. Industrial and Engineering Chemistry Process Designand Development 24, 852–858.

Skogestad, S., Jacobsen, E.W., Morari, M., 1990. Inadequacy of steady-state analysis of feedback control: distillate-bottom control of distillationcolumns. Industrial Engineering and Chemistry Research 29, 2339–2346.

Sneesby, M.G., Tade, M.O., Smith, T.N., 1998. Steady state transitions in thereactive distillation of MTBE. Computers and Chemical Engineering 22,879–892.

Sneesby, M.G., Tade, M.O., Smith, T.N., 1999. Two-point control of reactivedistillation column for composition and conversion. Journal of ProcessControl 9, 19–31.

Tian, Y.C., Zhou, F.T., Bisowarno, B.H., Tade, M.O., 2003. Pattern-basedpredictive control for ETBE reactive distillation. Journal of Process Control13, 57–67.

Vora, N., Daoutidis, P., 2001. Dynamics and control of an ethyl acetate reactivedistillation column. Industrial Engineering and Chemistry Research 40,833–849.

Wang, S.J., Wong, D.S.H., Lee, E.K., 2003a. Effect of interaction multiplicityon control system design for a MTBE reactive distillation column. Journalof Process Control 13, 503–515.

Wang, S.J., Wong, D.S.H., Lee, E.K., 2003b. Control of a reactive distillationcolumn in the kinetic regime for the synthesis of n-butyl acetate. IndustrialEngineering and Chemistry Research 42, 5182–5194.

towards further internal heat integration in design of reactive distillation

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