Control Strategy for Heat Integration in Heterogeneous Azeotropic Distillation

Control Strategy for Heat Integration in Heterogeneous Azeotropic Distillation


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Copyright © IFAC Low Cost Automation, Buenos Aires, Argentina, 1995


Brignole N.B., Tonelli S.M., Brignole E.A

PLAPIQUI - UNS - CONICET 12 de Octubre 1842 - 8000 Bahia Blanca Argentina FAX 054-91-565764

Abstract: The separation of azeotropic mixtures via heterogeneous distillation poses one of the most challenging process control problems. With the help of rigorous dynamic simulations, it has been possible to determine a suitable control strategy which enabled heat integration. First, new steady states for both columns were determined so as to allow operation under revamping conditions. Then, the energy recovery scheme was incorporated and tested. As no integration policy is feasible without proper regulation, the changes in the control policy were ascertained. Finally, the economic benefits were assessed by means of a cash flow analysis.

Keywords: Chemical industry - Control applications - Dynamic behaviour Ec.onomics Distillation columns - Integration - Mathematical models


Distillation has become common practice in industry in order to separate binary azeotropes. For the traditional system isopropanol water, the separation can be accomplished thanks to the addition of an entrainer, such as benzene or ciclohexane. Cases like this one are known as heterogeneous because the entrainer forms a ternary azeotrope, increases the relative volatilities between the key compounds and induces the formation of two liquid phases, which must be separated by decanting.

systems. The choice of the operating region significantly affects the possibility of effective operation and control. The presence of two liquid phases causes additional sensitivity because plant disturbances can induce undesirable inmiscibility inside the column. A careful choice of the thermodynamic model is required for accurate prediction. Numerical problems such as singular Jacobians and steady-state multiplicity frequently arise at the simulation stage.

Since distillation columns consume relatively large amounts of energy, it is worthwhile trying to reduce utility expenditures through heat integration, i.e. by making effective use of process streams for heating or cooling purposes.

In this work, the main goal is to achieve substantial energy savings both through heat integration and the determination of the most fav0l!rable operating conditions for two industrial distilt'ation columns,· simultaneously contemplating a significant increase in the plant's productive capacity. The control loops had to be defined with care because energy recovery schemes can be carried out in practice only when an adequate control strategy has been formulated and implemented.

Due to the complex aspects associated with heterogeneous azeotropic distillation in particular, this case study is remarkably interesting. A few papers have addressed the issue of controlling such 75

_____ ,l

C~ ::~.-.-_-_~~~~:J .---~--cIoo:J--I 0..... i i

. _.......... _1_


, -- - - - - - . - -- - - - - -


.. - -. ------ ---- -- -- ._-- .. _-.---- ----- - -· ·0 n.

Fig. 1. Flowsheet for the plant section under study. energy requirements for the heating and cooling devices.

2. PROBLEM STATEMENT This analysis was developed in the context of a technological-transfer project (Brignole et al, 1993) which also included a thorough examination of several other aspects such as the associated modifications to the existing flow streams, and the verification and design of the new pieces of equipment.

Table 1 specifies the steady state conditions determined for each column, operating with a 25% increase in the plant's original capacity. Table 1 Steady-state data for the azeotropic train under revamping conditions Variables (SI lDlits) Stages Pressure (bar) Feed tray (*) Feed flow (t) Feed composition (water / IPA) Feed temp./press Reflux ratio Distillate rate Distill. composition (water / IPA) Bottom flow Bottom comp. (waterlIPAlbenzene)

The train under study belongs to an existing plant which produces alcohols and ketones. This section in particular purifies iso-propyl alcohol (IPA). Figure 1 shows the corresponding flowsheet. The feed is an aqueous mixture with 80% weight water, ketones being its main impurity. The first column dehydrates it yielding a top product that should approach the composition of the binary azeotrope (88% weight alcohol), in order to get maximum recovery efficiency. The distillate enters the azeotropic column, where benzene is used as entrainer to break the binary azeotrope, hence making it possible to obtain pure IPA. As the ternary mixture fonns a low-boiling-point azeotrope, the top product approaches the azeotropic composition, whereas the bottom yields pure alcohol. The azeotrope can be split through condensation and cooling in two liquid phases. The aqueous heavy phase leaves the decanter while the light phase, rich in benzene, is reflux:ed back into the column. This property makes it remarkably easy to recover the entrainer. The makeup water fed to the decanter contributes to move the liquid-liquid equilibrium to mixtures with high contents of benzene. In this way, the minimum amount of water is recycled, thus reducing the

Top temp. Bottom temp. Steam flow Steam press Recycle flow Recycle comp. (water/ipa/benzene)

Dehydrating column 20 2.85 8 1676.0 0.918 0.082 388.0/3.0 1.45 206.0 0.3323 0.6677 2001.0 1.0 0.0 0.0 381.1 407.3 531.0 4

Makeup water Makeup benz. Decanter temp (0<) Tray I corresponds to the bonom tray. (t) All flows are expressed in kgmollh


Azeotropic column 25

1 17 206.0 0.3323 0.6677 340.0/1.0

125.0 0.000 0.995 0.005 338.8 359.7

372.7 0.100 0.228 0.772 160.0 1.8 308

The simulation policy adopted to reproduce plant behaviour and model the new conditions comprises several steps. Firstly, the reference steady states were adjusted to match already existing balances for the original design specifications. Then, operating conditions for the revamped process were established. This mainly involves choosing the reflux ratio for the first column and the amount of entrainer for the second one. This choice is difficult and the modelling of the posed problem is tricky due to the extreme sensitivity of the process.

3. PROCESS SIMULATION Most previous work on control of distillation columns has been focused on extremely simplified dynamic models and ideal mixtures. As those approaches lead to wrong profiles and ineffective controllers for azeotropic trains, powerful simulation tools become essential to ensure reliable results.

In this study, the preliminary steady-state values were found out by using GCDIST (Magnussen, 1982; Andersen, 1984). The rigorous dynamic simulations were obtained through the OYNSIM routines (Gani et al., 1992b), with the search of suitable initialization points for the state variables performed by means of steady-state simulations using SEPSIM ( Gani, 1990). Both the SEPSIM and OYNSIM packages with UNIF AC (Magnussen, 1982) as thermodynamic model were chosen because they had been tested successfully on problems which were similar to this case study (Bossen et al., 1993). OYNSIM contains specific equations for liquidliquid equilibrium and phase stability. Thermodynamic models play a very important role in azeotropic calculations because wrong phase diagrams can lead to erroneous design. Gani et al. (1992a) strongly recommend taking into account phase split, giving an example where the composition profiles change dramatically when the presence of two liquid phases has been neglected.

The impact of liquid phase splitting on the column was also analysed. The presence of two liquid phases on the trays should be avoided for it causes a significant decrease in efficiency. Therefore the azeotropic column was forced to operate at 338 K. This temperature is high enough to avoid this undesirable condition. In contrast, the decanter must have two liquid phases. To satisfy this requirement, the temperature was lowered to 308 K. 4. INTEGRATION POLICY The heat integration consists in capitalizing on the enthalpic contents of the top vapours leaving the dehydrating column to fulfill about 90% of the total energy requirements demanded by the azeotropic column. The eXistIng temperature difference between this stream and the vapour boilup of the second column was not big enough to allow the integration. This limitation was overcome by increasing the pressure in the first column from 1 bar to 2.85 bar. In this way, the temperature of the condensing vapours becomes 107 C, a satisfactory value for integration purposes.

The following assumptions were considered in the mathematical model: vapour holdups in trays and reboilers as well as heat losses to the surroundings are negligible; liquids are perfectly mixed and all trays are 100% efficient.

The increase in the thermal operating level of the dehydrating column changes significantly its energy requirements. Besides, its bottom product should leave this plant section at 106 C. Prior to the integration, the bottom had to be heated to satisfy this condition, whereas at 2.85 bar this product is so hot that it must be cooled instead. Thus, it was proposed to use the bottom to preheat the feed stream in 30 C, which is the amount needed so as not to cool the column internals.

The SEPSIM package contains Newton-Raphson techniques to solve the sets of algebraic equations. The methods make use of the system's Jacobians, which must be invertible matrices. This necessary condition is most often satisfied in conventional columns like the first one. In contrast, the azeotropic column exhibited almost singular Jacobians near the solution. As no alternative methods, such as those belonging to the homotopy-continuation family, were available, the problem was overcome by integrating the dynamic model without disturbances. In this way the initial estimates were brought close to the solution trajectory. It is interesting to note that these simulations must be short, otherwise the search induces entrainer's loss as impurity in the product streams and convergence is never achieved. This problem arises because the entrainer must be present inside the column at steady state although it is not fed to the column in significant quantities during continuous operation.


5.1 Checking the degrees offreedqm In order to avoid under- or over-specification of the control system, it is necessary to analyze the degrees of freedom by examining the system of differential and algebraic equations that represents the plant section . Under the assumptions stated in section 3, the non- linear dynamic model based on total and 77

the amount of water in the product stream, whereas increases cause the impurification of IFA with benzene. Figure 3 shows that the upper temperature sensor can monitor the amount of water, while the lower sensor detects IPA.

compound mass balances has E equations and V variables, calculated as follows: E = 3 (NA+ND) + 2 Ne (NA+ND) + 7 Ne + 16 V = 3 (NA+ND) + 2 Ne (NA+ND) + 11 Ne + 22

I. 0 ~_=_~--~--....__--..__-__, 0.12

where NA: number of trays in the azeotropic column, NO: number of trays in the dehydrating column, NC: number of compounds.

~:: t 0, 7

, "




~ '.



The difference V-E gives the degrees of freedom, which are reduced to 3 after having subtracted the input variables regarded as disturbances. The remaining degrees of freedom can be related to the main control objectives associated with the temperature loop in the dehydrating column, the bottom heat duty in the azeotropic column and the benzene-rich recycle from the decanter. Thus, the system is completely specified for the control scheme shown in figure 1.

_ 1.-353.73 lcInol/h • - . tJoJ65.48 "",ol/h - - L- 385. 92 kmol/h



0.4 1 0 .:1


O. 2

i_ _........_____ (1





19 20

Theoretical stages

Fig. 3. Molar fractions of IPA and water throughout the azeotropic column

5.2 Sensitivity analysis

5. 3 Control schemes

The dehydrating column normally operates at constant reflux. Simulations for different reflux values were carried out so as to study the effect on the temperature profile as well as the relationship between temperatures and the amount of organics in the stripping section. Operation at 1.25 reflux ratio is infeasible because the bottom product would contain more than 100 ppm organics. Figure 2 reveals a sensitive region for trays 4 to 6. As this variation corresponds to a change in the amount of organics, the sensor for the bottom-quality loop should be located there.

Control objectives and constraints: The great degree of sensitivity exhibited by azeotropic columns makes it common industrial practice to adopt a control policy of minimum deviation from steady state. Therefore, the global control objective is disturbance rejection.



,.....-,=:----------, ~


\ \

I 125 12(!

115 110 105

The original product-specification requirements naturally hold after revamping and integration. The isopropanol leaving the azeotropic column should always have less than 50 ppm and 5 ppm of water and benzene respectively. Besides, the total concentration of organics in the bottom-product stream withdrawn from the dehydrating column should be kept under 100 ppm in order to avoid product loss.



\ \


I ~


\ \

New control objectives are to be taken into account to ensure acceptable operation with heat integration. The main additional condition imposed on the first column is to maintain its top pressure in 2.85 bar so that the top vapours condensate around 107 C, thus yielding enough energy to heat the azeotropic column. Interactions caused by the energy recovery scheme should be minimized and satisfactory regulation must be achieved through the combined action of both reboilers. Avoiding the risk of flooding constitutes an extra requirement associated with operating feasibility.

\ R-1.45 \




lL.._ _ _ _ _ _ _ _ _ _ _ __







Theoretical stages

Fig. 2 Sensitivity analysis for the first column The temperature profiles for the azeotropic column exhibit two sensitive regions, at trays 3 to 5 and 16 to 18, where the temperature sensors can be placed. The column's extreme sensitivity induces operation at constant recycle. Small changes in the amount of entrainer ruin the product purity. Decreases in recycle rate (implying a reduction in the quantity of benzene entering the column) significantly augment

Loop by loop description: Revamping restricts the choice of controller structures, because the operating range of the prospective manipulated variables lies closer to the bounds imposed by the equipment


mixtures have big heats of vaporization (40,000 The simulation results for several disturbances confirmed the mild effect of pressure on condensation temperature.

limitations. All the control loops are schematized in Figure 1.


After evaluating several methods (Chin, 1979), it

was decided to regulate the condensation regime in

According to the RGA matrix, the loops do not interact; consequently, SI SO controllers can be designed.

the condenser/reboiler for pressure control by varying the condensing area with the valve at the condenser output. Using the live steam instead may be conflicting (Rademaker et al, 1975). Through simulations and l~sensitivity analysis it was concluded that the closed-loop operation of the reboiler/condenser always provides within 85% and 95% of the total energy without level saturation (0.4m - 1.3m) for pressures between 2.2 bar and 3.2 bar.

Dynamic considerations: a qualitative analysis of the dynamics associated with the main pairing (p - ml) was carried out. Although pressure loops respond very quickly for conventional columns, this result is not obvious in this case because the integration involves additional equipment. The influence of loop behaviour on the first column, as well as how heat exchanger dynamics affects performance should be taken into account. Assuming first order approximations and fast condensation dynamics a loop time constant of 2.8 min was calculated.

Since the specifications impose less than 100 ppm of organics in the bottom water, a quality loop that measures temperature at the most sensitive region and manipulates the injected vapour is indispensable. In order to minimize interaction with the pressure loop, the temperature measurement was corrected using the pressure values.

The new plant interconnections are the source of additional disturbances. Changes in the amount of energy supplied by the reboiler/condenser, for example, cause nonlinearities. The effect is shown in figure 4. A 5% step led to a substantial decrease in the plant's productive capacity, while a -5% variation ruined product purity.

The controllers for the azeotropic column aim at producing IPA containing less than 50 ppm water and 5 ppm benzene. The aforementioned condenser/reboiler yields around 85% of the heat duty. The already existing flooded-area reboiler provides the rest, thus ensuring quality and reducing the interactions caused by the integration. This loop monitors temperature in the bottom section and manipulates the condensate valve. As the integrated column works close to its flooding limit, it was necessary to incorporate a differential pressure controller that opens the bypass in case of overpressures.

0 ..


I::I 006




0~~~----~or,----~u2'----'o«~----~ol Otnl {t'l)

Fig. 4 Transient responses for the azeotropic column.

RGA analysis (Bristol, 1966): This tool was applied to the dehydrating column so as to measure interaction between the main pairings: pressurecondensate area (p-ml) and temperature-live steam (T-m2)· The calculation was based on the corresponding gain matrix, whose elements were obtained from simulations. As the mathematical models available work at constant pressure, the derivatives involving pressure variations were reformulated using the chain rule to give:

Flexibility considerations: a definite advantage of this proposal is the fact that the possibility of making the train work under the original conditions has been contemplated. Once the control scheme has been implemented, it is easy to switch to operation at atmospheric pressure in the dehydrating column. First, open bypass 1 and close valve 2. Then lower the bottom-temperature set point for the first column. In the azeotropic column, open the differential pressure loop and manipulate the original reboiler in full capacity.

!l- :1, :,1.



6. ECONOMIC PROFILE for i ,j = 1 , 2. To define the economic performance of the project a profitability analysis was carried out. The study was based on plant operation under revamping

The denominator, which comes from Clapeyron's equation, amounts to 105 K/bar because azeotropic 79

conditions, comparing profits with and without integration. Therefore, the proposed investment does not affect sales, while it yields savings in operating

The judicious use of existing resources is particularly important in developing countries. In this context, the energy recovering scheme becomes especially attractive because significant improvements in plant efficiency can be achieved by means of a low investment undertaking.


The additional pieces of equipment required for the integration are the reboiler/condenser, n.vo preheaters for the feed and reflux streams of the first column, a condensate pump and a flash tank. Based on their international purchase values, a total investment of 250,000 $ was estimated. This figure also takes into account the costs of the corresponding piping and accessories together with additional percentages for working capital, relocation, insulation, control and instrumentation. The expenditure makes it possible to get annual net energy savings of 110,000$. The value was calculated considering typical deduction rates for maintainance, overhead and taxes and assuming straight-line depreciation and construction completion within a year.

7. CONCLUSIONS The heat integration proposal and the analysis of the exact amount of entrainer led to significant savings. Its implementation was feasible thanks to an adequate control strategy. Although subjected to revamping restncttons, the proposed scheme satisfies the stringent quality requirements, allows energy integration minimizing the impact of disturbances, maintains the azeotropic column at work within ranges that guarantee condensation of the top vapours in n.vo liquid phases and has flexibility to operate under the original conditions. REFERENCES

It is interesting to note that the sole addition of the reboiler/condenser would not have made the

Andersen, P.M. (1984). Manual for constant molar overflow distillation programs: UNIDIST, GCDIST, SRKDIST. MAN 8402, Instituttet for Kemiteknik, Denmark. Bossen B.S., Jorgensen S.B. and Gani R (1993). Simulation, design and analysis of azeotropic distillation operations. Ind. Eng. Chem. Res., 32, 620-633. Brignole N.B., Cassino P., Colantonio M.C., Echarte R., Tonelli S.M., Urlic L. and Brignole E.A. (1993). Report 067/ 93, PLAPIQUl, Bahia Blanca, Argentina. Bristol E.H. (1966). On a new measure of interaction for multivariable process control. IEEE Trans. on Aut. Cont., AC-l1, 133-134. Chin T.G. (1979). Guide to distillation pressure control methods. Hydrocarbon ProceSSing, 10, 145-153. Gani R (1990) Manual for the steady state process simulator SEPSIM. Man 8702, Instituttet for Kemiteknik, Denmark. Gani Rand Fredenslund Aa. (1992a). Efficient and accurate computation of thermodynamic properties for design of separation processes. Trans 1 Chem E, Vol. 70, Part A, 7, 439-447. Gani R, Sorensen E.L. and Perregaard 1. (1992b). Design and analysis of chemical processes through DYNSIM. Ind. Eng. Chem. Res., 31, 1, 244-254. Magnussen T. (1982). Manual for UNIDIST: multi component distillation using UNIF AC. MA N 8102, Instituttet for Kemiteknik, Denmark. Rademaker 0 ., Rijnsdorp 1. and Maarveled A. (1975) . Dynamics and control of continuous distillation units. Elsevier, Oxford.

economics satisfactory enough. Due to its inherent operating savings, the integration between the feed and bottom streams of the dehydrating column turns out to be indispensable to ensure the viability of the project. eoo ,--_ _ _ _--,


r-P.I= 325 M






Fig. 5. Cash flow profiles for the addition of integration. Figure 4 illustrates the discounted and undiscounted cumulative cash flows. Curve I corresponds to zero discount rate. The actual interest on investment is reflected in curves 2 and 3 for estimations of the prevailing rates over the period in accordance with international and local rates of 7.5% and 14% respectively. The payback period (about 3.5 years) is within the project lifetime, estimated at 8 years. The net present values are positive, representing a gain on the project. The rates of return (TIR) corresponding to these cash flows have acceptable values because they are higher than the company's opportunity cost of money.