Multicomponent Distillation

Figure 1 Simple distillation } a still.

Multicomponent Distillation

V. Rico-RamOH rez and U. Diwekar,Carnegie Mellon University, Pittsburgh, PA, USACopyright^ 2000 Academic Press

Introduction

Distillation is the oldest separation process and themost widely used unit operation in industry. It in-volves the separation of a mixture based on the differ-ence in the boiling point (or volatility) of its compo-nents. The reason for the wide acceptance of distilla-tion is that, from both kinetic and thermodynamicpoints of view, distillation offers advantages overother existing processes for the separation of Suidmixtures:

1. Distillation has the potential for high mass trans-fer rates because, in general, in distillation thereare no inert materials or solids present.

2. The thermodynamic efRciency for distillation ishigher than the efRciency of most other availableprocesses in the chemical industry.

Designing a distillation column involves: (1) selectingthe type of column, mostly based on heuristics; (2)obtaining the vapour}liquid equilibrium data usingthermodynamics; and (3) Rnding the design variablessuch as number of equilibrium stages and operatingconditions required to obtain the desired separationbased on mass and energy balances.

When the mixture to be separated contains twocomponents, the design of a column can be accomp-lished by using graphical methods. However, formulticomponent systems the design methods aremore difRcult and are the focus of this article.

Fundamentals

Simple Distillation

Distillation began as a simple still. In such an opera-tion, a liquid mixture is heated (see Figure 1). Asa result, a vapour stream richer in the more volatilecomponents comes off, while the liquid, richer in theless volatile components, remains in the still. Thevapour stream is condensed and collected in the con-denser.

The analysis of simple distillation for a binarymixture presented in 1902 by Lord Rayleigh marksthe earliest theoretical work on distillation. ConsiderFigure 1. Let F (moles) be the initial feed to the

still and xF (mole fraction) be the composition ofcomponent A of the mixture. Let B be the number ofmoles of material remaining in the still, xB the molefraction of component A in the still, xD the molefraction of component A in the vapour dB producedduring an inRnitesimal time interval dt. The differen-tial material balance for component A can be writtenas:

ln BF"

xB

xF

dxBxD!xB

[1]

Complex mass and heat transfer processes occur indistillation processes and it is generally assumed thatthe vapour formed is in thermodynamic equilibriumwith the liquid. Hence, the vapour composition (xD)is related to the liquid composition (xB) by an equilib-rium relation of the functional form xD"f (xB).Note that, because of the unsteady nature of simpledistillation, the equilibrium relationship betweenxD and xB holds only for each inRnitesimal timeinterval dt.

The exact equilibrium relationship for a particularmixture may be obtained from a thermodynamicanalysis and is also dependent upon temperature andpressure.

Thermodynamics and Equilibrium Data

Accurate and reliable thermodynamic data forvapour}liquid equilibrium is essential to distillation

II /DISTILLATION /Multicomponent Distillation 1071

design. For binary mixtures, these data are generallypresented in the form of tables containing the liquidand vapour equilibrium compositions over a range oftemperatures for a Rxed pressure. The same informa-tion can also be plotted in what is called an x}ydiagram. For multicomponent mixtures, however,vapour liquid equilibrium data are difRcult to repres-ent in graphical or tabular form. In such case,K values are used instead.

K value and relative volatility The K value of acomponent i is a measure of the tendency of suchcomponent to vaporize. A K value is deRned by:

Ki"yixi

[2]

where yi is the equilibrium composition of the vapourphase for a composition xi of the liquid phase.K values are a function of temperature, pressure andcomposition, and they are widely reported for binaryandmulticomponentmixtures. An associated conceptis the relative volatility, i,j, which is a measure of theease of separation of components i and j by distilla-tion:

i,j"KiKj

[3]

Ideal and nonideal systems An ideal system is one inwhich the liquid phase obeys Raoults Law and thevapour phase obeys the ideal gas law. For such sys-tems, the K value is given by:

Ki"yixi"p

0i

P[4]

where p0i is the vapour pressure of pure componenti and P is the pressure of the system. Note that p0i isa function of temperature.

For a nonideal system, theK values can also dependupon the composition of the mixture and are ex-pressed in terms of fugacity coefRcients, where Vi isthe vapour phase fugacity coefRcient and Li is theliquid phase activity coefRcient, as given below:

Ki"LiVi

)p0iP

[5]

Azeotropic systems represent examples of nonidealmixtures for which eqn [5] has to be used.

Classi\cation of Distillation Processes

There are many criteria under which one can classifydistillation: type of accessories, operating mode,

design calculation assumptions, etc. Distillationcan either be binary or multicomponent. Accordingto the type of accessories used to increase the masstransfer in the separation process, a distillation col-umn can be packed (use of packing) or staged (use ofplates). It can be batch or continuous. Also, accordingto the assumptions made and accuracy expected ina distillation design calculation, a calculationtechnique can either be a shortcut method or arigorous method.

Packed columns and staged columns Althoughsimple distillation in a still historically represents thestart of the distillation process, a complete separationof the components of the mixture using this process isnot possible. Therefore, the application of these stillsis restricted to laboratory-scale distillation, wherehigh purities are not required or when the mixture iseasily separable.

One can look at simple distillation as consisting ofone equilibrium stage where a liquid and a vapour arein contact with one another and mass and heat trans-fers take place between the two phases. If N suchstages are stacked one above the other, and are al-lowed to have successive vaporization and condensa-tion, that results in a substantially richer vapour andweaker liquid (in terms of the more volatile compon-ent) in the condenser and the reboiler, respectively.This multistage arrangement is representative ofa distillation column, where the vapour from thereboiler rises to the top and the liquid from thecondenser is reSuxed downwards (see Figure 2). Thecontact between the liquid and the vapour phase isestablished through accessories such as packing orplates. When the accessory is a stack of plates, thenthe result is a column of trays. Similarly, if the acces-sory is packing, the result is a packed column.

Continuous distillation and batch distillation Thebasic difference between a batch column and acontinuous column is that in continuous distillationthe feed is continuously entering the column, while inbatch distillation the reboiler is normally fed at thebeginning of the operation. Also, while the top prod-ucts are removed continuously in both batch andcontinuous operations, there is no bottom product ina conventional batch distillation. Since in a continu-ous operation the total product Sow equals that ofincoming feed or feeds, the process reaches a steadystate. In batch distillation, on the other hand, thereboiler becomes depleted over time, so the processis unsteady. Such differences are illustrated inFigure 3.

Batch distillation is a direct extension of the simpledistillation still, where the Rayleigh equation

1072 II /DISTILLATION /Multicomponent Distillation

Figure 2 Equilibrium processes. (A) Single stage; (B) multi-stage.

Figure 3 Batch distillation versus continuous distillation.

(eqn [1]) is applicable. However, in both batch andcontinuous distillation, multistage mass transfer andthermodynamic equilibrium stage calculations areused for obtaining the steady-state relationship be-

tween the product composition (instantaneous in caseof batch) and feed composition.

Multicomponent MultistageEquilibrium Calculations

This section is divided in two parts. In the Rrst wediscuss approximate methods (or shortcut methods);the second part corresponds to rigorousmethods. Theapproaches are different depending upon the opera-tion mode of the column, that is, a continuous opera-tion or a batch operation.

In this section, our attention is focused on theapproaches to the design of continuous columns. Thereader can refer to the book by Diwekar (1995) forbatch distillation calculations.

Shortcut Methods

Approximate methods constitute a useful for the syn-thesis, analysis and design of distillation separations.The main advantage of shortcut methods is that theycan provide the feasible region of operation. Theyalso provide large saving in computer time, and some-times, they are sufRciently accurate that more expen-sive rigorous methods are not justiRed.

Concept of Nmin and Rmin Minimum number ofplates, Nmin, and minimum reSux, Rmin, are very im-portant concepts in the design of distillation pro-cesses, as they are considered to be the limiting condi-tions in the operation of a distillation column.


Nmin corresponds to the number of trays requiredfor separation in a situation in which the externalreSux ratio R (ratio of the liquid reSuxed to thedistillate rate) of the column is inRnite. This corres-ponds to total reSux operation.

Rmin corresponds to the minimum value of the ex-ternal reSux ratio required to achieve the speciRedseparation in a situation in which the number of traysof the column is inRnite.

Fenske}Underwood}Gilliland method The mostpopular of these shortcut methods is the Fenske-Underwood-Gilliland method (FUG). The basic as-sumptions of such a method are:

1. The system is ideal.2. Constantmolar overSow (as in theMcCabe Thiele

method for binary mixtures).3. The separation is essentially taking place between

the light key component and the heavy key com-ponent. The light key (lk) is the lightest compon-ent appearing in the bottom and the heavy key(hk) is the heaviest component appearing in thetop.

In the FUG method:

1. Fenskes equation is used to calculate the min-imum number of trays, Nmin.

2. Underwoods equation is used to estimate the min-imum reSux, Rmin.

3. Gillilands correlation is used to calculate the ac-tual number of trays, N (for any R given), or thereSux ratio, R, (for any N given) in terms ofprevious limiting values Nmin and Rmin.

The Fenske equation is:

Nmin"log

xDlkxBlk )

xBhkxDhk

log(lk,hk)[6]

where lk,hk is the relative volatility between the lightkey component and the heavy key component. Sinceit can be expected that the value of changes for eachtray of the column, the geometric average of thisvalue is generally used:

N"N ) N12 1 [7]

The Underwood equation can be written as:

i

i ) xi,Di!

"Rmin#1 [8]

where is a root of the equation:

i

i ) xi,Fi!

"1!q [9]

such that hk44lk. hk and lk are the relativevolatilities of the key components (light and heavy) inthe calculation. As stated earlier, such componentsare the ones that the designer uses as the basis for theseparation.

Finally, the Gilliland correlation is given by:

N!NminN#1 "1!exp

1#54.4G11#117.2G )

G!1G0.5

[10]

where

G"R!RminR#1 [11]

The main assumptions of the Underwood equationare the assumption of constant molar Sow rates andan ideal system. Such assumptions constitute themain limitation of the algorithm.

Rigorous Methods

Recent developments in computer hardware and soft-ware have made it possible to use rigorous methodsfor the design of distillation processes. In thesemethods, the assumption of constant molar Sow ratesis no longer considered. The implication of removingsuch an assumption is that rigorous methods not onlyconsider mass balances, but also enthalpy balancesfor each of the trays of the column. Thus, rigorousmethods require simultaneous convergence of massand energy equations. Depending on the calculationsequence, there are several rigorousmethods reportedin the literature. The most important of thesemethods are: (1) Thiele}Geddes; (2) tridiagonalmethods; (3) Naphtali}Sandholm; (4) inside-out algo-rithms; (5) convergence methods; and (6) 2N New-tonmethods. The method of Naphtali}Sandholm andthe inside-out algorithm, which are commonly usednowadays, are discussed in this work to give an ideaof the scope and applications of rigorous methods.

MESH equations Most rigorous methods involvethe solution of the so-called MESH equations. Foreach stage n in a distillation column (and for eachcomponent i in a mixture of C components), theequations representing mass balance (M), equilib-rium relationships (E), summation of compositions(S) and energy balance (H), constitute the MESHequations. In addition, both K values and enthalpies


Figure 4 Equilibrium stage. Derivation of MESH equations.

are generally given as functions of temperatures, pres-sures and compositions. The generalized form of theMESH equations for the equilibrium stage shown inFigure 4 and the expressions for K values and enthal-pies are present in Table 1.

Naphtali}Sandholm method In the Naphtali}Sandholm method, the number of variables of theMESH equations is reduced by the introduction ofcomponent Sow rates and side streams. Furthermore,the summation of compositions are eliminated. ThosemodiRcations result in the equations presented inTable 2.

To solve the system of MESH equations given inTable 2, the vectors of variables and equations areordered as follows. Variables:

XM "[XM 1, XM 2,2, XM n,2, XM N] [12]

where N is the number of stages and

XM n"[vn,1, vn,2,2, vn,C, Tn, ln,1, ln,2,2, ln,C]T [13]

Equations:

FM "[FM 1, FM 2,2, FM n,2, FM N] [14]

where

FM n"[HK n, Mn,1, Mn,2,2, Mn,C, En,1, En,2,2, En,C]T

[15]

The solution process is iterative, using one of theseveral variations of the Newton method. Thus, cor-rections at each iteration k are obtained from


Table 1 MESH equations

Relationship Equation

Mass balance Ln#1 ) xn#1,i#Vn1 ) yn1,i#Fn ) zn,i!(Ln#Un) ) xn,i!(Vn#Wn) ) yn,i"0Equilibrium yn,i"Kn,i ) xn,iSummation of compositions

iyn,i!1"0

H energy balance Ln#1 ) hn#1#Vn1 ) Hn1#Fn ) HFn!(Ln#Un) ) hn!(Vn#Wn) ) Hn!Qn"0

K values and enthalpies Kn,i"Kn,i (Tn, Pn, xnyn)Hn,i"Hn,i (Tn , Pn, yn)hn,i"hn,i (Tn, Pn, xn)

Table 2 MEH equations for method of Naphtali and Sandholm

Relationship Equation

Component flow ratesand side streams

vn,i"yn,i ) Vn

ln,i"xn,i ) Lnfn,i"zn,i ) Fnsn"Un/LnSn"Wn/Vn

M Mn,i"ln,i ) (1#sn)#vn,i ) (1#Sn)!ln#1,i!vn1,i!fn,i

E En,i"Kn,i ) ln,i ) k vn,k/k ln,k!vn,i"0H

HK n"hn ) (1#sj) ) i

ln,i

#Hn ) (1#Sn) ) i

vn,i

!hn#1 ) i

ln#1,i!Hn1 ) ivn1,i

!HFn ) i

fn,i!Qn"0

(classical Newton}Raphson equations):

FM (k)"!FMXM

1

(k)

) FM (k) [16]

XM (k#1)"XM (k)#t ) XM (k) [17]

where t is such that 04t41. t is the factor thatensures progress toward the solution of the system atequations of each iteration.

Inside-out algorithm In the Naphtali}Sandholmmethod, the temperatures and component Sowratesare the primary solution variables (see eqn [13]) and

are used to generate the K values and enthalpies fromcomplex correlations. Hence, such a method updatesthe primary variables in an outer loop, with theK values and enthalpies updated in an inner loopwhenever the primary variables change.

In inside-out algorithms, the previous situation isreversed. These methods use complex K values andenthalpy correlations to generate parameters forsimple K values and enthalpy models. Hence, theseparameters become the variables for the outside loop.The inside loop then consists of the MESH equations.In every step through the outside loop, the simpleK values and enthalpy models are updated by usingtheMESH variables from the inside loop. This sets upthe next pass through the inside loop. The book byKister (1992) provides detailed guidelines for the useof the various inside-out methods.

Special Separations

When the components of a mixture have low relativevolatilities, or when the mixture contains a largenumber of components, separation by distillation be-comes difRcult and expensive because a large numberof trays or a large number of columns are required forthe separation. Furthermore, some systems may shownonideal behaviour such as the formation of azeo-tropes or a reversal of the relative volatility with thechange in pressure from top to bottom in a column.Complex systems which have these characteristics arecommon in the pharmaceutical and synthetic chem-ical industry.

This section presents a brief review of separationsin which the traditional distillation process is altered,but the general principles of multicomponent distilla-tion still apply. Three broad categories of such specialseparations exist: azeotropic distillation, extractivedistillation and reactive distillation. Petroleum distil-lation will also be discussed since it represents a case


Figure 5 Azeotropic behaviour.

in which the complexity of the mixture (petroleum)requires special considerations for the separation.

Azeotropic Distillation

Highly nonideal systems, with components havingclose boiling points among them, often produce azeo-tropes. Azeotropes can be identiRed by using an x}ydiagram. When an azeotrope is present, the equilib-rium curve crosses the line x"y (453 line), as shownin Figure 5.

Azeotropes limit the separation that can beachieved by conventional distillation. Sometimes it ispossible to shift the equilibrium by changing the pres-sure of the system sufRciently to move the azeotropeaway from the region where the separation must bemade. Other cases, however, require the addition ofa new material in order to achieve separation.

In azeotropic distillation, the equilibrium behav-iour of the mixture is modiRed by adding a newmaterial (called the solvent or entrainer). The addedentrainer forms a minimum boiling point azeotropewith one or more components and distils overhead.The distillate is generally heterogeneous, that is, it iscomposed of two immiscible liquids when condensed.Such a heterogeneous nature facilitates the separationof the product from the entrainer.

Extractive Distillation

Extractive distillation also involves the addition ofthe third component to the mixture (solvent or en-trainer). However, in the case of extractive distilla-tion, the solvent is a relatively high boiling pointmaterial, which is present at high concentration oneach stage and exits at the bottom. To improve theefRciency of the process, the entrainer has to be added

at the top of the column, so that its concentration oneach stage will be enough to produce the desiredeffect in the equilibrium of the original mixture. Fi-nally, the entrainer is separated from the bottomsproduct in another distillation column.

Reactive Distillation

The idea of combining reaction and separation ina single apparatus has been extensively investigated.Doherty and Buzad (1992) present a survey of theavailable design techniques for reactive distillation.Reactive distillation is particularly attractive when-ever a chemical reaction provides the favourable ef-fect of reacting away azeotropic mixtures so that thebehaviour of the liquid phase is simpliRed. In addi-tion, it has been shown that reactive distillationhas the potential of eliminating recycle costs whena liquid reaction involves a large excess of onereactant.

In general, the current trend in reactive distillationdesign is using experimental results from bench-scaleproblems in the initial stages of the design, and thenusing computer-aided simulation tools for scale-upand operability issues.

Possible proRtable applications of reactive distilla-tion processes are numerous. However, an incom-plete understanding of the interactions of the manynonlinear phenomena such as chemical reaction,phase equilibrium, mass transfer and countercurrentSow has prevented the widespread use of such pro-cesses. Considerable research effort in the area iscurrently being conducted.

Petroleum Distillation

Petroleum distillation is particularly difRcult becauseof the large number of components of the mixtureand large scale of the processes. This type of distilla-tion involves products that are not easily identiRablecomponents. Instead, separation is achieved in termsof pseudo-components, which are generally charac-terized in terms of their true boiling point ranges(TBP), an average relative molecular mass and an APIgravity. TBP data are widely available and are gener-ally presented in form of curves.

There are two main approaches to the design ofpetroleum distillation columns. The Rrst consists ofthe solution of mass and energy balances based onempirical correlations, and is basically a calculationby hand. This approach was developed by Packie.

In the second approach, each pseudo-component ischaracterized for properties (such as vapour pressureand enthalpy) by using homologous-series ap-proaches. Thus, rigorous mass and energy balancescan then be applied to determine the separation in


terms of the reSux ratio. Several efRcient computerprograms following this approach have been de-veloped.

Packed Columns

Several approaches exist for the design of packedcolumns. These are based on the concepts of numberof transfer units (NTU), height of transfer units(HTU) and height equivalent to a theoretical plate(HETP). The last of these concepts is the most widelyused.

Since methods for the design of staged distillationcolumns are well developed, a common approach isto calculate the number of trays N using such ap-proaches and then to Rnd the height of the packedcolumn, h, by the relation:

h"N ) HETP [18]

There exist various correlations for predicting thevalue of the HETP. One of most commonly used isthe Sherwood correlation. It can be expected thatHETP will change with respect to the operating con-ditions, physical properties of the liquid, etc., so, it iscalculated in terms of correlations containing manyfactors.

Nonequilibrium Distillation

All the mathematical methods (binary, rigorous,shortcut) presented earlier assume that each stage inthe column is an equilibrium stage. In reality, how-ever, this assumption is rarely satisRed.

Stage Ef\ciency

An approach to nonequilibrium calculations is theuse of the concept of stage efRciency. The most com-mon approach is to modify the rigorousmethods withthe introduction of the so-called MurphreeefRciency in the calculations. The Murphree efRcien-cy in a stage calculation can be deRned as:

ELMi"xout,i!xin,ixi!xin,i

[19]

for the liquid and

EVMi"yout,i!xin,iyi!yin,i

[20]

for the vapour. xi are the compositions of the liquidthat would be in equilibriumwith the outlet composi-tion of the vapour. yi are the compositions of the

liquid that would be in equilibrium with the outletcomposition of the liquid.

Mass Transfer Rates

It has been shown that stage efRciency prediction andscale-up are difRcult and unreliable. For highlynonideal, polar and reactive systems, a transport phe-nomena approach for predicting mass transfer rates ispreferred. Such mass transfer rates are calculatedcontinuously along the column similarly to the HETPcalculation for packed columns.

Nonequilibriummodels for the calculation of masstransfer rates assume that, while the bulk vapour andliquid phase are not in equilibrium with each other,there is an equilibrium at the interface. Hence,the net loss or gain for a component at the interfaceis expressed in a rate form. For instance, the netgain by the vapour because of the transfer at theinterface is:

NV0

ij "Nvij ) daj [21]

whereNVij is the vapour Sux of the component at somepoint through the interface and daj is the interfacearea through which the Sux passes. The mass transferrates for liquid and vapour,NVij andNLij, are dependenton the mass transfer coefRcients for each phase. Thereexist several correlations for the heat and mass trans-fer coefRcients and these are dependent on the com-positions in the bulk phase, the temperatures in thebulk phase and interface, and on the packing or traygeometries.

Industrial Applications

Distillation is by far the most widely used separationtechnique in the petroleum, natural gas and chemicalindustries so, applications of multicomponent distil-lation are numerous. A couple of industrial applica-tions are described in this section.

Primary Distillation of Crude Oil

A typical conRguration for the distillation of a crudeoil unit includes two main columns, an atmospherictower and a vacuum tower (see Figure 6). In theatmospheric tower, crude oil is rectiRed (at a pressureno greater than 275.8 kPa (40 psi); to yield a distillateproduct containing light hydrocarbon gas, light andheavy naphtha, kerosene, diesel oil, and a bottomproduct of heavier components (TBP greater than4203C). Each of the side streams of the atmospherictower are sent to side strippers that have a partialreboiler or steam stripper. The side stream strippersserve to remove the light components. Stripping by


Figure 6 Crude oil distillation unit.

steam is also frequently used in the bottom of thetower.

The bottom product of the atmospheric tower isfurther separated by rectiRcation in the vacuumtower. The feed-tray pressure of a vacuum tower isusually 6 kPa (45 Torr). Vacuum towers are mainlydesigned to obtain heavy distillates such as gas oil,

lubricating oils and bunker fuels with asphalt as thebottom product.

The pump-around systems shown in both of thetowers serve to make much larger liquid Sows on theintermediate stages and produce a net increase inliquid Sow. This serves as a point of control to keepthe plates from running dry.


Figure 7 Separation of products of the manufacture of ethylene and propylene.

Highly developed procedures for the preliminarydesign of fractionators that process petroleum arecommercially available through computer programs.The program REFINE of the ChemShare Corpora-tion and the PROCESS (now PRO-II) program ofSimulation Sciences Inc. are two examples.

Ethylene and Propylene Production

The manufacture of ethylene and propylene is one ofthe most important operations of the petrochemicalindustry. In that process, ethylene and propylene areformed from the thermal cracking of other hydrocar-bons, such as ethane, propane and naphtha. Themixture resulting from the thermal cracking is verycomplex. Hence, the mixture has to be separated intorelatively pure ethylene and propylene, ethane andpropane to be used as a recycle, methane and hydro-gen to be used as fuel, and heavier products to be usedfor gasoline. A typical reRnery gas feed to the separ-ation system of this process contains hydrogen,ethylene, methane, ethane, propane, propyleneand lower compositions of other heavy hydro-carbons. The distillation sequence most commonlyused for the separation of the mixture is shown inFigure 7.

In a high pressure plant (no refrigeration is neededfor condensation of products), the distillation se-quence consists of Rve distillation columns:

1. Demethanizer2. Deethanizer3. Ethylene/ethane separator4. Depropanizer5. Propylene/propane separator.

Both the propylene/propane and the ethylene/ethane separator require high towers with large dia-meters because such mixtures contain componentswith very close relative volatilities. A plant that usesthe conRguration described here was built by PullmanKellogg Inc., Houston, Texas.

In the case of a lower pressure plant, the deethan-izer precedes the demethanizer because refrigerationis required for the feed of the demethanizer. So, byplacing the deethanizer Rrst, important utility savingsare obtained.

Future Work

Enormous progress has been made on the applicationand design of distillation technology. However, chal-lenges still exist in some areas, which lead to thefollowing ongoing research:

1. Improvement of mass transfer coefRcients inpacked distillation columns. Great effort is beingmade on the design of efRcient packings and accu-rate correlation of their performance.


2. The simulation, synthesis and design of reactiveand azeotropic distillation. Such topics still consti-tute a gap in the knowledge of distillation tech-nology.

3. Investigation of complex conRgurations for batchdistillation processes.

4. Use of optimization methods for obtaining opti-mal conRguration and design of batch and con-tinuous distillation processes.

5. Online optimization and control of columns.

See also: II/Distillation: Batch Distillation; Theory of Dis-tillation; Vapour-Liquid Equilibrium: Correlation and Pre-diction; Vapour-Liquid Equilibrium: Theory.

Further Reading

Diwekar UM (1995) Batch Distillation: Simulation, Opti-mal Design and Control. Series in Chemical and Mech-anical Engineering. Washington, DC: Taylor & Francis.

Doherty MF and Buzad G (1992) Reactive distillation bydesign. Transactions of the Institution of Chemical En-gineers 70: part A.

Gmehling J and Onken U (1977) Vapor}Liquid Equilib-rium Data Collections, DECHEMA Chemistry Dataseries, vol. 1. Frankfurt:

Henley EJ and Seader JD (1981) Equilibrium-Stage Separ-ation Operations in Chemical Engineering. New York:Wiley.

Holland CD (1981) Fundamentals of Multicomponent Dis-tillation. New York: McGraw-Hill.

King CJ (1980) Separation Processes, 2nd edn. New York:McGraw-Hill.

Kister HZ (1992) Distillation Design. New York:McGraw-Hill.

Perry RH, Green DW and Maloney JO (1984) PerrysChemical Engineers Handbook, 6th edn. New York:McGraw-Hill.

Schweitzer PA (1979) Handbook of Separation Techniquesfor Chemical Engineers. New York: McGraw-Hill, TheKingsport Press.

Treybal RE (1980) Mass Transfer Operations, 3rd edn.New York: McGraw-Hill.

Packed Columns: Design and Performance

L. Klemas, Bogota, ColombiaJ. A. Bonilla, Ellicott City, MD, USACopyright^ 2000 Academic Press

Use of Packing in Distillation

Use of packing in mass transfer has its origins in theearly 1800s for simple applications such as alcoholdistillation, and in sulfuric acid plant absorbers. Glassballs, coke or even stones were used as packing ma-terials. Nevertheless packings for distillation were notestablished until the 1930s with the use of regularshape materials such as ceramic Raschig rings andBerl saddles, as well as the availability of distillationcalculations such as the McCabe}Thiele and Pon-chon}Savarit methods. Early in the second half of thecentury, the use of packing for distillation wentthrough a transformation, producing the second-generation packings (see Table 1). Regular and im-proved shape of packings, such as pall rings, becameavailable with larger open areas that permitted a sub-stantial increase both in capacity and column efRcien-cy. In the 1960s Sulzer introduced the wire-meshpackings with very high efRciency (low height equiva-lent to a theoretical plate, HETP), resulting in a newtransformation in the use of packings. In the 1970s

and 1980s all major mass-transfer equipment manu-facturers developed structured packings. Comparedto the traditional tray columns spectacular improve-ments in plant capacity were achieved, but also someprojects were pitfalls, when the expected beneRts didnot materialize. Manufacturers started realizing thatliquid distributors had to be improved, but there wasno coherent understanding, nor correlations, thatcould lead to a safe distributor-column system design.Many manufacturers returned to trays, producingnew improved designs, using the area under thedowncomer for vapour Sow: these trays are offeredwith new names that indicate their increased vapourSow capacity (MaxySow, Superfrack, etc.). The needfor good distribution and its effect on the columnefRciency are now well understood, allowing safedesign and efRcient applications for random andstructured packings in large industrial columns.

General Concepts

Distillation separation is based in relative volatilitythat makes it possible to concentrate the more volatilecomponents in the vapour phase while the less vol-atile ones remain in the liquid phase. Distillationcolumns are countercurrent vapour}liquid mass-transfer devices, where the required separation andpuriRcation of components is achieved.

II /DISTILLATION /Packed Columns: Design and Performance 1081

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Multicomponent Distillation