Intermediate Heat Exchange for Fixed Separation Requirements: Applications to a Binary Sieve Tray...

INTERMEDIATE HEAT EXCHANGE FOR FIXED SEPARATIONREQUIREMENTS

Applications to a Binary Sieve Tray Distillation Column forEnergy Savings

I. NIKLASSON BJORN1, U. GREN2� and A. P. SOEMARDJI2

1AstraZeneca R&D Molndal, Molndal, Sweden2Chalmers University of Technology, Department of Chemical Engineering & Environmental Science, Gothenburg, Sweden

The aim of this study is to investigate different ways of introducing intermediate heatexchange into a given sieve tray distillation column by using computer simulation.The number of stages and the product qualities were kept constant before and after

the introduction of the intermediate heat exchange. The intermediate heat exchange wasachieved by applying an open heat pump circuit where column streams were withdrawnfrom the column and after heat exchange, they were returned to the column immediately.Different return positions were evaluated as well as arrangements for adding/removingheat to/from different positions in the column were studied. The rate of the side streamswas also varied. The column has been considered as a stand-alone system.

The intermediate heat exchange was analysed from different points of view. It was foundthat the entropy production could be decreased in a tray column if the side streams were rein-troduced in an optimal way. A 7% reduction in entropy production was obtained for thecolumn studied. A decrease in external energy of about 14% has been calculated with thereturn of the optimal side streams into the intermediate heat exchange. With the same modi-fications, a 27% decrease in the variance of the driving forces was achieved, which indicatesthat the system has become more reversible in its nature.

The liquid and the vapour flows change along the column due to the intermediate heatingand cooling. This will of course affect the mixing on the trays and the residence time, andconsequently, the Murphree tray efficiencies will change. An average value of the Murphreeand Chan-Fair efficiency for the optimal case was slightly higher compared to that for theconventional case.

Keywords: diabetic distillation; intermediate heat exchange; entropy production; energysavings.

INTRODUCTION

Distillation is the most common separation process inindustry. The chemical process industries, including thepetroleum and chemical industries, account for about27% of the energy demand in the United States. Separationprocesses used to recover and purify products account forover 40% of this energy demand (Humphrey et al., 1991).Distillation is a process, which uses considerable

amounts of energy. This is because heat is used as the

main separating agent (see e.g., King, 1980; Seader andHenley, 2006). Therefore it is of vital importance that aproper analysis is made regarding the placement of the dis-tillation operation within the process plant and how it isintegrated (see e.g., Smith, 2005; Seider et al., 2004).There are also interesting possibilities of combiningdistillation columns in the separation of multi-componentmixtures. To improve the energetic situation for the caseof a binary distillation and a stand alone column, manymethods have been discussed in the literature (see e.g.,Freshwater, 1951; Mah et al., 1977; Naka et al., 1980;King, 1980; Takamatsu et al., 1988) and the concept of dia-batic or quasi-reversible distillation (see e.g., Tondeur andKvaalen, 1987; Rivero et al., 1994; LeGoff et al., 1996) isone of these. A classification of different possibilities for

�Correspondence to: Professor U. Gren, Chalmers University of Technol-ogy, Department of Chemical Engineering & Environmental Science,SE-41296 Gothenburg, Sweden.E-mail: [email protected]

453

0263–8762/06/$30.00+0.00# 2006 Institution of Chemical Engineers

www.icheme.org/cherd Trans IChemE, Part A, June 2006doi: 10.1205/cherd.05154 Chemical Engineering Research and Design, 84(A6): 453–464

improvement thermodynamically of the distillation oper-ation has been presented by Fonyo and Mizsey (1994)and Klemes and Stehlik (2005).Under diabatic conditions, the distribution of heat and

consequently the driving forces, are changed along thecolumn by the use of intermediate heat exchange. Changingthe conditions to those of the limiting case of a reversiblecolumn would lead to an infinite number of trays, whichobviously would not be an economical or practical alterna-tive to the ordinary configuration. However, for a columnwith fixed number of trays it will be necessary to increasethe gross heating and cooling duties in order to fulfil thebase case purity requirements. Consequently, it would bebeneficial to use heat pump connections for this purpose.Evaluation of diabatic distillation can be made using

different measures. One of these measures has beenbased on the energy demand and the distribution of theintermediate heating and cooling (Fonyo, 1973; Fonyoand Foldes, 1974a, b; Naka et al., 1980; Kayihan, 1980;Dhole and Linnhoff, 1993; Agrawal and Herron, 1998a,b; Nakaiwa et al., 2001). In Agrawal and Herron (1998b)e.g., heuristics have been presented to judge if anintermediate reboiler or an intermediate condenser will beeffective in improving the thermodynamic efficiency ofthe distillation process or not. Depending on the feedproperties the methodology will answer which one of thetwo alternatives will be the most effective one.In Nakaiwa et al. (2003) a review on heat-integrated

distillation columns has been presented. The quantificationof the energy demand is of great importance in many indus-trial applications since it is directly connected to operatingcosts.In Rivero et al. (1994) the exergy effectiveness was used

as an evaluation parameter of diabatic distillation where theeffective conversion from thermal exergy to chemicalexergy was dealt with. The results showed a maximum inexergy effectiveness as an effect of the minimization inentropy. For a diabatic rectifying column it was foundthat the exergy destroyed in a diabatic column is 0.5times less than for the ordinary (adiabatic) column. Theamount of coolant that was used in the reboiler was reducedby a factor 14 in the diabatic column compared to theadiabatic one.The minimum work of separation and its relation to

entropy efficiency for the process operation was discussedin Fonyo (1973). Theoretically, the minimum work of sep-aration can be achieved for a totally reversible distillation.This will however require an infinite number of trays.When examining improvement possibilities for a basecase distillation column the entropy efficiency will not bea relevant measure if the separation requirements are keptconstant, as only the entropy production will change.Another measure for evaluation is the entropy production

of the system (Mullins and Berry, 1984; Tondeur andKvaalen, 1987; De Koejier and Kjelstrup, 2000; Schalleret al., 2001; Bjorn et al., 2002; Rosjorde and Kjelstrup,2005) and the entropy production in the whole systemcould be minimised by proper distribution.The isoforce principle was discussed in Sauar et al.

(1996). In distillation minimum entropy production rate issaid to be obtained when the driving forces are uniformlydistributed over the column. In De Koejier et al. (2002)the isoforce concept was applied when adding two

intermediate heat exchangers to a column. A rule ofthumb was proposed: The entropy production in thecolumn should be distributed proportionally to the variationof the phenomenological coefficients, related to heat ofvaporization. For the case studied this meant that theheat exchangers should be places directly around the feedtray. As diabatic distillation requires extra heat exchangersthe methodology was further developed in De Koejieret al. (2004) where the effect of heat exchange distributionon minimum entropy rate was studied.

As a basis for the present study a column was usedpresented in Bjorn et al. (2002, 2006). The experimentalsieve tray pilot plant distillation system has a rather flexiblestructure. The column system has been constructed to allowintermediate heating and cooling to be exchanged along thecolumn in an open heat pump circuit with an external heatexchanger. This means that intermediate heat exchange canbe studied for several side stream connections in thestripping section and in the rectifying section. Theexperimental column presented in Naito et al. (2000) isan interesting realisation of the so-called heat integrateddistillation column (HIDiC) (Takamatsu et al., 1988),where the stripping section and the rectifying section areseparated into two columns. The heat integration isachieved by heat exchangers internally connected betweenthe columns. However the concept restricts the number ofheat exchange possibilities and the flexibility in flowdistribution.

The aim of the study was to investigate the practicalapplication of introducing intermediate heating and coolinginto a given tray column by withdrawing side streams to anopen heat pump circuit for energy savings. Especially thedifference to a situation where the number of trays couldbe allowed to vary freely was of particular interest in thisstudy. The consequences of the side stream return methodand positions were considered using different measures.Examples of measures were energy demand, entropy pro-duction and variance. The variance of the driving forceswas considered as a practical method of evaluating theequipartitioning of the driving forces for the differentcases studied. Furthermore, the influence of the introduc-tion of intermediate heating and cooling on the tray effi-ciency was taken into account. The number of trays andthe purity of the products were kept constant as comparedto conventional distillation.

THEORY

In order to evaluate and compare alternatives for ener-getic and/or thermodynamic improvement of a base caseseparation the following parameters (measures) and sometheoretical backgrounds will be presented below.

External Energy Demand

When introducing intermediate heat exchange into anexisting column with defined quality requirements andconstant number of trays, the gross heating and coolingwill have to increase. In this case, the introduction ofintermediate heat exchange can be defended only ifthe decrease of external energy demand is larger thanthe extra need for circulated energy. The saving inexternal energy for this non-conventional open heat pump

Trans IChemE, Part A, Chemical Engineering Research and Design, 2006, 84(A6): 453–464

454 NIKLASSON et al.

connection is calculated as:

EES% ¼QR,conv � (QR þ 3 � W)

QR,conv

� �� 100 (1)

The factor 3 was chosen as to make adjustment for therelation between the cost of electricity and the cost ofadmission steam in accordance with that used in severalcountries (Moser and Schnitzer, 1985).

Entropy Changes in a Separation Process

In a separation process the entropy change can beexpressed as:

DSprocess ¼ DSproduction � DSconsumption (2)

The production term is related to the heat transportbetween the reboiler and the condenser, and the consump-tion term is introduced because of the separation of com-ponents. If the separation was reversible, the productionand consumption terms would be equal (Fonyo, 1973).One frequently used measure of the reversibility is theentropy efficiency as expressed by Fonyo as:

h ¼DSconsumption

DSproduction

(3)

Entropy Consumption

The change in entropy for the component separation inthe minimum isothermal conditions (King, 1980) can beexpressed as:

DSsep¼�F �Rgas

(XD�XB)�

�(XF�XB)

� XD �logXF

XD

þ(1�XD)�log(1�XF)

(1�XD)

� �þ(XD�XF)

� XB �logXF

XB

þ(1�XB)�log(1�XF)

(1�XB)

� ��(4)

In the expression the assumption is made that the productshave the same temperature and pressure as the feed. For thespecial case, where the number of trays and the purities ofthe products are kept constant as compared to conventionaldistillation the entropy consumption is constant

Entropy Production

In the present study, a change in entropy production is ofdecisive importance to the evaluation of the simulationresults. This has to do with the fact that the defined entropyconsumption will not change when there is no change inseparation task.In a conventional distillation column the degradation of

heat can be written as:

DSprod,conv ¼QC

TC

�QR

TR

(5)

When intermediate heat exchange points are introduced,the entropy production equation for the system indicated incan be written as:

DSprod,ihex ¼QC

TC

�QR

TR

þQi

Ti

�Qj

Tj

(6)

Variance

Based on thermodynamics the best design would be tohave equal driving forces. The driving forces on say eachtray could be based on e.g., the chemical potential, but tosimplify it could be expressed as a concentration differencedue to the major influence of the mass transfer compared toheat transfer and mixing effects (e.g., De Koeijer, 2002). Inthe present work we chose to take the difference betweenthe vapour phase, y, and the liquid phase concentrationexpressed as the equivalent to the equilibrium concen-tration, y�, as the driving force. Ideally this driving forceshould be constant along the column. Taking the variance,s2 as it is normally defined (Perry et al., 1997) by

s 2 ¼½N � S(y� � y)2 � (S(y� � y))2�

(N � (N � 1))(7)

as a measure of the deviation from this ideal situation it willgive an idea of how well adopted the column is to an idealdiabatic distillation.

Mass Transfer Considerations

A change in the internal flows caused by the introductionof intermediate heat exchange in an existing column will ofcourse influence the mixing on the trays, and consequentlycause changes in the mass transfer efficiencies on the trays.One of the common methods that can be used to predict theefficiencies is the (AIChE, 1958).

In the AIChE method, the tray point efficiency is calcu-lated from the number of overall gas phase transfer units as:

EOG ¼ 1� e�NOG (8)

This relation is derived from the assumption of a totallymixed liquid on any vertical line through the liquid on thetray. It is also assumed that the gas phase passes the tray inplug flow. If total mixing were achieved both vertically andhorizontally in the direction of the liquid flow, then yout,Eqwould be the same over the whole tray when xout is thesame, so that EOG ¼ EMV.

In reality, total mixing is never fulfilled. The Pecletnumber, Pe, specifies the degree of back mixing in theliquid phase on a tray. A value of zero indicates completemixing, and with a Peclet number approaching infinity, aplug flow describes the flow pattern. The AIChE methodconsiders this by using a relation between the pointefficiency and the Murphree tray efficiency, namely

EMV

EOG

¼1� e�(hþPe)

(hþ Pe) � 1þ ð(hþ Pe)=h½ Þ

þeh � 1

h � 1þ ðh=hþ Pe½ Þ�(9)


INTERMEDIATE HEAT EXCHANGE FOR FIXED SEPARATION REQUIREMENTS 455

where

h ¼Pe

2� 1þ

4 � L � EOG

Pe

� �1=2

�1

!(10)

SIMULATION PROCEDURES

The simulations were performed using a computer steadystate simulation program called PRO II 5.11 fromSimulations Sciences. In order to include for e.g,. tray effi-ciencies in the simulations these were calculated using acalculator routine based on the physical dimensions of thecolumn, ingtray size, active area and flow rates on the trays.The column used in these studies was a real fixed-

distillation column with 12 sieve trays (see Figure 1).The column diameter is 0.2 m and the total height is 6 m.The column is equipped with a reboiler and a totalcondenser.1 The equipment was designed to separateethanol from n-propanol at atmospheric pressure.The separation requirements were based on the introduc-

tion into the mid-part of the column of a defined mixture ofethanol and n-propanol at a certain feed rate. The purityof the distillate and bottom products was kept constant.The basic equilibrium model is non-random two liquid(NRTL). The algorithm for the solution is inside out.The first step of the simulations was to find out how to

reintroduce side streams after external heat exchange inan open heat pump circuit. From different ways of reintro-ducing side streams, the best one is found and tried withdifferent locations of intermediate heat exchange (sidestreams) in the column. For this step the feed flow ratewas 1.5 kmol h21, feed composition XF ¼ 0.5, top and bot-toms flow rates 0.75 kmol h21. Top and bottoms productpurities were XD ¼ 0.9 and XB ¼ 0.1 at normal pressure.The rectifying section is the upper part of the column,

from the feed-tray above. Therefore, in this section, somevapour was withdrawn and returned to the column afterthe phase changes in order to make changes on the operat-ing line. At the bottom part of the column, the strippingsection, the process is the reversed; the liquid is withdrawnand returned as vapour to the column. These changes in theoperating line were intended in order to make possible anequal distribution of the driving force.Three different ways of reintroduction of side streams

were studied, namely: Simulation A: Return to a trayabove in the rectifying section and to a tray below in thestripping section. Simulation B: Return to the same positionof the removal for both sections. Simulation C: Return tothe tray below in the rectifying section and the trayabove in the stripping section (conf. Lynd and Grehtlein,1986).The aim of the second step was to identify the best

locations of the side streams. As can be seen in Table 1,there are several possible points of reintroduction of theside streams into the rectification and stripping sefctionsof the column. The alternatives were first studied separatelyand then the best locations from each section were com-bined and simulated together in the complex system. Forthe system studied, the side reintroduction of the streams

was chosen so that phase and composition coincided ascompared to the base case.

For this second step a higher feed flow rate,100 kmol h21 was chosen to be more close to industrialpractice. The feed composition XF was 40 mole-% ethanol.Top and bottoms product purities were XD ¼ 0.92 andXB ¼ 0.05 at normal pressure.

It could be noted that in the heat exchangers the duty wasset by the operating conditions i.e., flow rate and a phasechange of 95% minimum. The minimum temperaturedifference of the heat exchanger was normally chosen tobe 3 K.

RESULTS AND DISCUSSION

Side Stream Return Position Arrangements

The objective of introducing intermediate heat exchangeis to improve the efficiency and quality of the separation,when the number of trays is fixed. From the left chart ofFigure 2 it is evident that, for the specified separationtask, it is obvious that the gross energy demand, heredefined as the total heating duties added to the system

Table 1. Possible locations of intermediate heatexchange in the column.

Rectifying section Stripping section

4 ! 2 9 ! 115 ! 3 10 ! 126 ! 4 11 ! 137 ! 5

Figure 1. The sieve tray column.

1The trays in the simulation program are numbered 1 to 14, with 1 as thecondenser and 14 as the reboiler.



(in the reboiler and in the intermediate heat exchanger), has toincrease when the intermediate heat exchange is introduced.It may be noted that there are differences between

methods used when returning the side streams, and thedifferences become most drastic for Simulation C. This isdue to the large amounts of extra energy that have to beexchanged in the reboiler and condenser to compensatefor the re-mixing in connection with improper side streamreturn and constraints to guarantee product purity. For theoptimal case, i.e., Simulation A, the gross energy demandrises only slightly with increasing side stream flux.However, a distinction should be made between the gross

energy demand and the external energy demand which isdefined as the energy demand for a system consisting ofa reboiler and an open heat pump, see equation (1). A sig-nificant reduction in the latter is obtained in the case ofoptimal return, see the right hand chart in Figure 2. Thisreduction occurs even if the gross energy demand isincreased. The savings in external energy were found tobe as high as 14% in the optimal case, i.e., SimulationA. They were noticeable also when the side stream wasreturned to the same tray, i.e., in Simulation B, whereasthere was a negative energy saving i.e., a demand of exter-nal energy in the non-optimal case, i.e., Simulation C.Taking the entropy production as a measure of the ther-

modynamic quality, Figure 3 indicates that an optimal sidestream withdrawal rate can be obtained. For the columnstudied, it coincides with a total flow rate correspondingin size to that of the feed rate. This optimal side streammethod together with the side stream ratio of 0.5 gives areduction in entropy production of 7%. If the side streamsare returned as proposed in Simulations B and C, however,no optimal side stream rate is obtained. It may therefore beconcluded that, in order to increase the thermodynamic effi-ciency, it is of primary importance that the side streams bereturned in an optimal way to avoid mixing streams, whichdiffer in composition and therefore counteracts the separ-ation. By the flexibility with different positions for withdra-wal and return in Simulation A the extra operating lineswill allow for a better alignment with the equilibrium line.For a certain side stream rate, the energy demands Qi and

Qj and the temperatures are practically the same for allmethods of side stream return, which makes the

contributions of the terms Qi/Ti2 Qj/Tj in equation (6)approximately the same. To perform the same separationin all methods, the extra energy needed in the condenserand reboiler in Simulations B and C must overcome theincrease in entropy production due to mixing within thecolumn. The increase in entropy due to the mixing ofstreams of different compositions could be represented bythe same type of equation as DSsep in equation (4).

A minimum in entropy production may be seen as aresult of a balance between the mixing entropy productionand the decrease in entropy production due to the locationof the duties at more appropriate temperatures. With such aproper balance, it is possible to obtain a minimum inentropy production even for a case where the side streamis returned to the point of withdrawal, if the energyexchange at temperatures close to the feed gives largeenough contributions to the entropy production.

The influence of the side streams on the reflux ratio wasalso investigated. In some of the simulations, the refluxratio was found to increase in some and to decreasein some, compared to the ordinary distillation as inFigure 4, the left chart. This is due to the changes invapour flow to the condenser keeping the distillate rate

Figure 2. Gross energy demand (left) and external energy savings (right) versus relative side stream flow rate for different side stream return methods.

Figure 3. Entropy production versus relative side stream flow rate fordifferent side stream return methods.



constant. In Simulation C, the reflux ratio was increasinginstead of decreasing. This was because the returningstream composition was not matched with the tray compo-sition and the re-mixing on the return tray had to becompensated by an increase in the reflux ratio.A comparison using the variance as a measure showed

slightly different results. Simulations A and B were foundto give improved distribution of the driving forces as com-pared to the base case. However, in Simulation C, the var-iance was increased in all cases, see Figure 4, the rightchart. This means that the distribution of the drivingforces to the system was not favourable. In other wordsthe slope of the operating lines changed extremely so theside streams did not improve the process. In Simulations,A and B, the changes of the slopes were more favourableresulting in an optimal operating situation.A relevant question when discussing the possibility of

side streams and intermediate heat exchange is whetherthe proposed combinations of flows in the column influencethe performance in such a way as to counteract the ambi-tions of improving the operation. Calculations were per-formed in order to determine the Murphree tray efficiencyand the operating limits. In Figure 5 variations of the

Murphree tray efficiency along the column are shown forthe different simulation models studied. It can be notedthat the marked change in Emv is strongly related to thechange in vapour flow rates as discussed in (Bjorn et al.,2002, 2006). It may be concluded that, for the sidestream rates presented and for the case A where phaseand composition coincide after side stream return, theMurphree tray efficiency is improved in the top andbottom of the column. It may also be concluded that theMurphree tray efficiency varies for the different sidestream rates, see Figure 6. Simulation A has a larger aver-age value of the Murphree tray efficiency in the cases withside stream to feed stream ratio of 0.5–1.0, which coincideswith those cases that have a lower entropy production thanthe conventional case. In the case of the non-optimal returnmethod, Simulation C, the efficiency is decreased every-where in the column, and the worst situation arises forthe largest side stream rate where the largest mixing ofstreams with different compositions occurs.

To summarize the first part it is quite obvious that afterwithdrawal and heat exchange of a side stream it is of theoutmost importance to return the phase changed streamso that it will support the overall separation task for the

Figure 4. Reflux ratio (left) and variance (right) versus relative side stream flow rate for different side stream return methods.

Figure 5. Murphree vapour tray efficiency for different side stream returnmethods versus PROII tray number. Figure 6. Mean Murphree vapour tray efficiency in the column versus.



system. The best results were found when the returned heatexchanged streams, no matter vapour or liquid, werereturned to positions were phase and composition coincideas closely as possible.With a proper side stream return it is possible to achieve

external energy savings, decrease in entropy production,more evenly distributed driving forces represented by a mini-mum in the variance and an increase in the average Murphreetray efficiency. This can be achieved with a fixed number oftrays and the same separation as in the base case.

Location of Intermediate Heat ExchangerArrangements

The return of the side stream from heat exchangers canof course be applied to several locations in a column. Inthis study the objective is to find the best locations oftwo interconnected heat exchanger arrangements. Thelocation of an intermediate heat exchanger in the columnmeans the tray position of withdrawing and returning thestreams. In the rectifying section, the liquid was returnedto a position above the tray of withdrawal and vice versafor the stripping section where the vapour was returned toa position below the position of the withdrawal. The possi-bilities to withdraw vapour/liquid and return liquid/vapourin each section are shown in Table 1.As a basis in the search for advantageous locations of the

intermediate heat exchangers, composition differences arepresented for the base case in Figure 7. These compositiondifferences can be taken as measures of the driving forces

along the column. In an ideal system, thermodynamically,these driving forces should be constant.

The resulting curves for the pilot plant column show thedriving forces for three different feed concentrations. It canbe concluded that there is an observable variation in thedriving force along the column. It can be concluded thatthe stripping part can be readily improved. In the rectifyingpart, however, the improvements may be less pronounced.It can also be concluded that the skew will make a correctbalance between the rectifying and stripping part difficult.In order to illustrate the effect of introducing intermediateheat exchange the case with a feed concentration of 40%ethanol was chosen. The best location in each section ofintroducing intermediate heat exchange arrangement willbe discussed as well as results of using an interconnectedsystem.

Introducing intermediate heat exchanger arrangementin the rectifying section

There are some possibilities of combinations to withdrawvapour and return liquid in the rectifying section, these are4 ! 2, 5 ! 3, 6 ! 4 and 7 ! 5. Since the side stream tofeed flow ratio may have a small influence on the result, thesimulations covered three different ratios: 0.4, 0.5 and 0.6.The results can be seen in Table 2.

When the vapour side stream is taken out in the rectify-ing part of the column less vapour is going up to the con-denser and the condenser duty will decrease. At the sametime the system will keep the distillate rate constant byreducing the amount of liquid returned. This reductionwill reduce the reflux ratio (L/D) and the slope of the oper-ating line (L/V). The results show that introducing a sidestream into the rectifying section gives reductions in thereflux ratio and condenser duty but gives an increase inreboiler duty and variance. The increase in reboiler dutyserves to compensate for the gross increase in coolingduty. The increase in gross cooling/heating duties is aresult of that the number of trays and the separationrequirements are kept constant. The best location seemsto be 4 ! 2 which also holds for still lower side streamto feed flow ratios.

As discussed earlier the possibilities for improving thevariance by a single side stream intermediate heat exchan-ger in the rectifying section are limited. This is due to theexistence of evenly distributed driving forces already inthe base case. The higher up the side stream is withdrawn,less liquid will be returned and this makes the slope lesssteep and the line’s point of interception with the Y-axisis raised, closer to the operating line. Increasing the sidestream to feed flow ratio will not improve the results, seeFigure 8.

Figure 7. The driving force along the column expressed as vapourcomposition differences.

Table 2. Different heat exchanger arrangements and flow ratios in the rectifying section.

4 ! 2 5 ! 3 6 ! 4 7 ! 5

Base 0.4 0.5 0.6 0.4 0.5 0.6 0.4 0.5 0.6 0.4 0.5 0.6

Reflux Ratio 2.6361 1.7192 1.4905 1.2594 1.7512 1.5365 1.3272 1.8164 1.6293 1.4483 1.9124 1.7541 1.6065s 2 . 104 2.5821 3.2162 4.0726 4.3015 3.3044 3.6107 3.9607 3.4938 3.9288 4.4566 3.6945 4.2491 5.5611Cond. duty, Gcal/h 1.3847 1.0347 0.9477 0.8598 1.0469 0.9652 0.8855 1.0717 1.0005 0.9316 1.1082 1.0480 0.9918Reb. duty, Gcal/h 1.3974 1.4298 1.4383 1.4483 1.4433 1.4574 1.4734 1.4694 1.4942 1.5212 1.5071 1.5431 1.5832HE, Gcal/h 0 0.383 0.478 0.573 0.384 0.479 0.575 0.385 0.481 0.577 0.386 0.482 0.579



Introducing an intermediate heat exchanger arrangementin the stripping sectionThe possibilities to withdraw liquid and return vapour in

the stripping section are 9 ! 11, 10 ! 12 and 11 ! 13.The results can be seen in Figure 9. The side stream tofeed flow ratio simulations covered three different ratios:0.4, 0.5 and 0.6.The results show that introducing the side stream into the

stripping section gives reductions in the variance and reboi-ler duty but gives an increase in reflux ratio and condenserduty. The introduction of a side stream in the stripping partwill help levelling out the large variations in driving forcesfor the base case. This will result in a reduction of thevariance as shown in Figure 9. Introducing a side streamin the stripping section is resulting in similar effects aswhen introduced in the rectifying section, see Figure 10.

Introducing an intermediate heat exchanger arrangementinto the interconnected systemIn order to integrate the duty requirements in the strip-

ping and rectifying parts of the column, two side streamarrangements were used—the interconnected system—asshown in Figure 11. The interconnected system is anopen heat pump system and requires that the amount of

side stream used in the stripping and the rectifying sectionsshould be balanced with respect to the latent heattransferred.

In the simulations of the interconnected system, locationsand flow rates were chosen based on the results in the var-iance of the stripping section (see the circle on Figure 9).The pair in the stripping section that gives the smallest var-iance is 9 ! 11. This pair will be selected for the intercon-nected system. A side stream to feed flow ratio of 0.6 willbe used.

In Table 3 results are shown for different possible pairlocations in the rectifying sections connected to the alreadychosen one in the stripper section. The side stream to feedflow ratio of 0.6. It is clear from the results that for the usedside stream to feed flow ratio the lowest possible varianceof the driving forces can be found for the 9–11 and 4–2.

Figure 8. Operating line for the rectifying section and connection 5 ! 3.

Figure 10. Operating line for the stripping section and the connection10 ! 12.

Figure 11. Flowchart of an interconnected system.Figure 9. Influence of side stream location and side stream flow ratios inthe stripping section.



The effect of the side streams on the operating lines canbe illustrated with a McCabe Thiele diagram as shownin Figure 12. The operating lines for the interconnectedsystems were found to give more smooth lines and to bebetter aligned with the equilibrium curve. The upperposition of the side stream will, however, create a bumpin the upper operating line.This could be explained clearly if the driving force along

the column is plotted for the simulations as in Figure 13. Inthis figure, the distribution of the driving force for pairs4 ! 2 and 9 ! 11 is almost constant along the columnand a jump is only obtained in the top part of the column

(where the x is high). This corresponds to the lowest var-iance obtained for this combination. It should be notedthat with the configuration and restrictions given for thisstudy there has to be a balance of the duty between thestripper and the rectifying section. This will evidentlyforce the system to converge differently to a situationwhere the duty was distributed optimally with respect toeach section. Thus, the optimum location was not theone, which could be anticipated, based on an analysis ofeach section.

General comments on the introduction of anintermediate heat exchanger arrangement in theinterconnected system

For the selected interconnected system the resultsobtained based on variance were compared with othermeasures of performance and for different side stream tofeed flow ratios. It can be concluded that the varianceand the DSproduction covariate. This is, however, not surpris-ing as considering their thermodynamic relationship. How-ever, the minimum occurs at slightly different side streamflow rates, showing a minimum in entropy production forthe side stream to feed ratio at 0.6 and for the variancefor 0.4.

In Figure 14 it is shown how the potential for a decreasein entropy production varies with side stream to feed flowratio for different feed composition (top and bottom pro-duct compositions kept constant). It is clear from the resultsthat for an application with a higher feed composition toseparate it is possible to achieve a larger decrease inentropy production than for a lower feed composition.

Table 3. Influence of rectifying section path for the interconnected system. Side stream to feed flow ratio 0.6.

Base 7 ! 5 6 ! 4 5 ! 3 4 ! 2

Ethanol comp. in bottom 0.05 0.05 0.05 0.05 0.05Reflux ratio 2.6361 1.8111 1.6928 1.5792 1.4984s 2 . 104 2.5821 2.5706 2.3442 2.0493 1.8320Cond. duty, Gcal/h 1.3847 1.0720 1.0246 0.9814 0.9506Reb. duty, Gcal/h 1.3974 1.0847 1.0374 0.9941 0.9633

Figure 12. The operating lines for the interconnected system with different locations of the rectifying side stream path. The side stream to feed flow ratio is 0.6.

Figure 13. The driving force profile along the column for theinterconnected system with different locations of rectifying side stream path.



In this case the decrease is 14% for feed composition 60%and as low as 5% if the feed composition is 40%. A largerfeed composition in the base case indicates a larger poten-tial for improvements when studying the driving forceprofiles of the base cases, see Figure 7. It can also be con-cluded from the base cases that higher driving forces allowfor larger heat exchange. So for the highest feed compo-sition, 60%, the minimum in entropy production is foundat a higher side stream to feed flow ratio then for 40%.A more even distribution of the driving forces along the

column known from the base cases for a lower feed concen-tration will also give the lowest variances. A similar trend(as shown in Figures 7 and 14) is obtained for the varianceas shown in Figure 15.In Figure 16 a comparison is made between different

locations of interconnected intermediate heat exchangersfor the feed composition 60%. The figure shows a

considerable decrease in entropy production for connec-tions 5–3 and 7–9 of 16%.

Regarding the gross energy demand there is a steadyincrease with increasing side stream flow ratio as expectedbut the EES% is also increasing, however subdued withincreasing side stream flow rate. However, no obviousoptima were found but the maximum is set by the operatingconstraints of the column.

Since the column studied in this work is rather short astudy was also made with a column with the doublenumber of trays. The results show a similar trend for themeasures used. However, it can also be concluded thatfurther improvements of the thermodynamic performancecould be made. It was found that including two pairs ofside stream units will make it possible to better align theperformance with the best one from a thermodynamicpoint of view. The units will also better balance aroundthe feed position.

Comments on different measures and economicalconsequences of intermediate heat exchange

When analysing the influence of different intermediateheat exchange methods the result is to a large extent influ-enced by the evaluation methods used. Especially from anindustrial perspective the economical consequences areimportant. In order to make a just comparison to industrialpractice a slightly larger column (23 trays and a diameter of0.85 m) and feed flow (100 kmol h21) with an open heatpump was chosen. As can be seen in Table 4 (Davies,2004) a strong relationship between the external energysavings, EES%, measure and the economic analysisexpressed as return on investment, ROI, or the presentvalue. The investments considered are those that areadded due to the instalment of the intermediate heatexchange. The greatest saving was observed for the casewhen the side streams were located at the extremes of thecolumn. In comparison with the thermodynamically moreattractive outline as shown by e.g., the variance this discre-pancy may be attributed to the impact of the energy cost.Interestingly, it was also found that introducing multiplesets of side streams a much better agreement between thedifferent methods was found.

These findings would thus suggest that when consideringwhich diabatic distillation configuration should be adopted,the external energy savings play a great role on theeconomic aspects of the column.

Figure 15. Influence of feed composition on the variance productionprofile for different side stream to feed flow ratio for connections 4–2and 9–11.

Figure 16. Influence of different locations of the interconnected intermedi-ate heat exchanger arrangements versus side stream to feed flow ratio, feedcomposition 60%.

Figure 14. Influence of feed composition on the entropy production fordifferent side stream to feed flow ratio for connections 4–2 and 9–11.

Table 4. Examples of economical measures of different side streamconfigurations. Feed rate 100 kmol h21, distillate rate 35 kmol h21,xF ¼ 0.35, xD ¼ 0.90.

Configuration ROI %Present value,

SEK Variance . 104 EES%

5–3 14–18 108 5.58 1.23 22.35–3 17–20 168 6.49 1.45 25.25–3 21–23 185 6.96 2.56 279–6 14–18 90 4.83 0.8 19.19–6 17–20 143 5.66 1.15 21.79–6 21–23 164 6.24 2.42 23.812–8 14–18 53 3.47 0.69 14.112–8 17–20 79 4.29 1.21 16.712–8 21–23 87 4.73 2.59 18.35–3 21–23 8–12 14–18 94 10.12 0.77 40.1



CONCLUSIONS

The aim of this study is to investigate different ways ofintroducing intermediate heat exchange into a given sievetray distillation column by using computer simulation.The number of stages and the product qualities were keptconstant before and after the introduction of the intermedi-ate heat exchange. It should also be observed that energeticimprovements always have to be achieved with a minimumof extra equipment. Our results yield the followingconclusions:

. The restriction to use a column where the number oftrays is fixed in contrast to let the number of traysincrease without limits will render situations where theenergy savings will turn negative and the entropy pro-duction will increase as compared to the conventional case.

. A substantial decrease in external energy demand isobtainable by the use of intermediate heating and coolingtogether with an open heat pump circuit, when the separ-ation requirements and the number of trays are heldconstant.

. It is essential that the streams are returned correctly tothe column after the heat exchange has been performed.Non-optimally reintroduced streams will result in anincrease in the external energy requirements.

. The entropy production can be decreased in a trayed dis-tillation column by using intermediate heat exchange,with the product requirements and number of trayskept constant as in the conventional case.

. The changes in internal flows tend to give an increasedaverage Murphree tray efficiency along the column forthe optimal case, whereas it decreases for the non-optimal cases. The Murphree tray efficiency alsoincreases toward the ends of the column as a result ofthe change in mass transfer.

The chosen sequence depends on the aim and conditionof the process or the priority. For a fixed column and theseparation of different components the variance or entropyproduction seems to be the best measure. For the aim ofreducing the energy involved it seems better to useenergy demand or EES% as a measure than variance aslong as the same separation achieved.

NOMENCLATURE

DE eddy diffusivityEES% external energy savingsEOG point efficiencyEMV Murphree tray efficiencyF feed flow rateh entropy efficiency, variable in equations (9) and (10)L ratio of the slopes of the equilibrium and operating linesN number of stages, including condenser and reboiler, in this

case 14NOG number of overall mass transfer units for the vapour phasePe Peclet number, ZL

2/DE. tL

Q heat or dutyRgas gas constantROI return on investmentDS entropy changetL residence time for the liquid on the plateT temperatureW compressor powerX liquid mole fractionY and y vapour mole fraction

y� equilibrium vapour mol fractionZL length of liquid paths 2 variance of the driving force

SubscriptsB bottomsC condenserD distillateEq equilibrium conditionsF feedi heat exchange point in enriching sectionj heat exchange point in stripping sectionR reboiler

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The manuscript was received 11 July 2005 and accepted for publicationafter revision 7 March 2006.



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Intermediate Heat Exchange for Fixed Separation Requirements: Applications to a Binary Sieve Tray...

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