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Practical EngineeringGuidelines for Processing

Plant SolutionsFeb 2007

Author:

Chew Yin Hoon

Distillation Column Selection andSizing

(ENGINEERING DESIGN GUIDELINES)

Checked by:

Karl Kolmetz

This design guideline are believed to be as accurate as possible, but are very general not for specific designcases. They were designed for engineers to do preliminary designs and process specification sheets. The finaldesign must always be guaranteed for the service selected by the manufacturing vendor, but these guidelines willgreatly reduce the amount of up front engineering hours that are required to develop the final design. Theguidelines are a training tool for young engineers or a resource for engineers with experience.

This document is entrusted to the recipient personally, but the copyright remains with us. It must not be copied,reproduced or in any way communicated or made accessible to third parties without our written consent.

TABLE OF CONTENTS

INTRODUCTION

I) Distillation History 4

II) Types of Distillation Processes 4

III) Mode of Operation 5

IV) Column Internals 5

Scope 8

General Design Consideration 9

The Selection of Column Internals 11

DEFINITIONS 14

NOMENCLATURES 15

THEORY OF DISTILLATION COLUMN DESIGN 17

A) Vapor-Liquid Equilibrium (VLE) 17

I) Ideal behavior in both phases 17

II) Liquid phase non-idealities 18

Local composition equations 18

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Practical EngineeringGuidelines for Processing Plant

Solutions

Distillation Column Selection and Sizing

(ENGINEERING DESIGN GUIDELINES)Feb 2007

This design guideline are believed to be as accurate as possible, but are very general not for specific design cases. Theywere designed for engineers to do preliminary designs and process specification sheets. The final design must always beguaranteed for the service selected by the manufacturing vendor, but these guidelines will greatly reduce the amount of upfront engineering hours that are required to develop the final design. The guidelines are a training tool for young engineersor a resource for engineers with experience.


Activity coefficient estimation methods 19

III) High-pressure systems 19

Vapor fugacity coefficient, liquid activity coefficient 20

Vapor and liquid fugacity coefficients 21

B) Phase Distribution 21

I) Incompressible liquid 22

II) Negligible Poynting correction 22

III) Vapor obeying the ideal gas law 22

IV) Ideal liquid solution 22

V) Other methods to determine K-values 23

24C) Distillation Calculations

24I) Product specifications

24II) Column operating pressure

26III) Minimum number of stages

26IV) Minimum reflux ratio

27V) Operating reflux ratio and number of stages

28VI) Tower diameter and height

APPLICATION 29

REFEREENCES 30

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SPECIFICATION DATA SHEET 31

CALCULATION SPREADSHEET 32

LIST OF TABLE

Table 1: Pressure drop in difference services 13Table 2: Common temperature differences for difference types of medium

in condenser and Reboiler. 25

LIST OF FIGURE

Figure 1: Schematic Diagram of Distillation Column/ Fractionator. 6

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INTRODUCTION

Distillation is by far the most important separation process in the petroleum and chemicalindustries. It is the separation of key components in a mixture by the difference in theirrelative volatility, or boiling points. It is also known as fractional distillation or fractionation.

In most cases, distillation is the most economical separating method for liquid mixtures.However, it can be energy intensive. Distillation can consume more than 50% of a plantsoperating energy cost. There are alternatives to distillation process such as solventextraction, membrane separation or adsorption process. On the other hand, theseprocesses often have higher investment costs. Therefore, distillation remains the mainchoice in the industry, especially in large-scale applications.

I) Distillation History

The history of distillation dated back to centuries ago. The full history of distillation hasbeen chronicled by Forbes in 19481. Reputedly, it was the Chinese who discovered itduring the middle of the Chou dynasty. It was later introduced to India, Arabia, Britain andthe rest of the world.

Early distillation consisted of simple batch stills to produce ethanol. Crude ethanol wasplaced in a still and heated, and the vapor drawn from the still was condensed forconsumption. Lamp oil was later produced using the same method, with crude oil heatedin batch stills.

The next progression in the history of distillation was to continually feed the still andrecover the light product. Further advancements include placing the stills in series andinterchanging the vapor and liquid from each still to improve recovery. This was the firsttype of counter-current distillation column that we have today.

II) Types of Distillation Processes

There are many types of distillation processes. Each type has its own characteristics andis designed to perform specific types of separations. These variations appear due todifficulty in separation when the physical properties of the components in a mixture arevery close to one another, such as an azeotropic mixture.

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One type of variation of the distillation processes is extractive distillation. In this type ofprocess, an external solvent is added to the system to increase the separation. Theexternal solvent changes the relative volatility between two close components byextracting one of the components, forming a ternary mixture with different properties. Thesolvent is recycled into the system after the extracted component is separated from it.

A distillation column may also have a catalyst bed and reaction occurring in it. This type ofcolumn is called a reactive distillation column. The targeted component reacts when it isin contact with the catalyst, thereby separated from the rest of the components in themixture.

III) Mode of Operation

Distillation towers can be classified into two main categories, based on their mode ofoperation. The two classes are batch distillation and continuous distillation.

In batch distillation, the feed to the column is introduced batch-wise. The column is firstcharged with a batch and then the distillation process is carried out. When the desiredtask is achieved, the next batch of feed is introduced. Batch distillation is usually preferredin the pharmaceutical industries and for the production of seasonal products.

On the other hand, continuous distillation handles a continuous feed stream. Nointerruption occurs during the operation of a continuous distillation column unless there isa problem with the column or surrounding unit operations. Continuous columns arecapable of handling high throughputs. Besides, additional variations can be utilized in acontinuous distillation column, such as multiple feed points and multiple product drawingpoints. Therefore, continuous columns are the more common of the two modes,especially in the petroleum and chemical industries.

IV) Column Internals

Column internals are being equipped into distillation columns to provide better mass and

heat transfers between the liquid and vapor phases in the column. These include trays,packings, distributors and redistributors, baffles and etc. They promote an intimatecontact between both phases. The type of internals selected would determine the heightand diameter of a column for a specified duty because different designs have variouscapacities and efficiencies. The two main types of column internals discussed in thisguideline are trays and packing.

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There are many types of trays or plates, such as sieve, bubble-cap and valve trays.Packing, on the other hand, can be categorized into random and structured packing. Inrandom packing, rings and saddles are dumped into the column randomly whilestructured packing is stacked in a regular pattern in the column.

Figure 1: Schematic Diagram of Distillation Column/ Fractionator.

Figure 1 shows a schematic diagram of an example distillation column or fractionator. Thefeed enters the column as liquid, vapor or a mixture of vapor-liquid. The vapor phase thattravels up the column is in contact with the liquid phase that travels down. The process

above the feed tray is known as rectification (where the vapor phase is continuallyenriched in the light components which will finally make up the overhead product) whilethe process below the feed tray is known as stripping (as the heavier components arebeing stripped off and concentrated in the liquid phase to form the bottom product). At thetop of the column, vapor enters the condenser where heat is removed. Some liquid isreturned to the column as reflux to limit the loss of heavy components overhead.

S

Feed

Reboiler

OverheadReceiver

Condenser

Bottom

Overhead

Product

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At each separation stage (each tray or a theoretical stage in the packing), the vaporenters from the stage below at a higher temperature while the liquid stream enters fromthe stage above at a lower temperature. Heat and mass transfer occur such that theexiting streams (bubble point liquid and dew point vapor at the same temperature andpressure) are in equilibrium with each other.

The condenser above the column can be either a total or partial condenser. In a totalcondenser, all vapor leaving the top of the column is condensed to liquid so that the refluxstream and overhead product have the same composition. In a partial condenser, only aportion of the vapor entering the condenser is condensed to liquid. In most cases, thecondensed liquid is refluxed into the column and the overhead product drawn is in thevapor form. On the other hand, there are some cases where only part of the condensedliquid is refluxed. In these cases, there will be two overhead products, one a liquid withthe same composition as the reflux stream while the other is a vapor product that is inequilibrium with the liquid reflux.

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INTRODUCTION

Scope

This design guideline covers the basic element in designing a typical distillation column,which includes column selection and sizing.

In designing a distillation column, the thermodynamics of the vapor and liquid phasesmust be understood. The vapor-liquid equilibrium (VLE) determines the minimum numberof stages required to achieve the degree of separation needed. The minimum reflux ratioalso depends on the VLE data of the mixture.

A few equations that are commonly used in the industry are illustrated in this guideline toestimate the minimum number of stages and the minimum reflux ratio of a column basedon the VLE data, such as the Fenske-Underwood equation. Some design heuristics arealso highlighted. These rules are based on design experiences and take into account boththe safety and economical factors.

The selection of column internals is very critical in distillation column design. There is awide variety of trays and packing in the market. Each design has its strengths andweaknesses. However, the quotations from vendors are sometimes contradictory andconfusing. This could lead to a wrong choice of column internal. Therefore, some generalconsiderations are depicted to aid engineers in making the right choice of columninternals.

A distillation column is sized by determining the diameter and height of the tower. Aninitial estimation of the tower diameter can be done based on the vapor and liquidloadings in the column. The height of the tower, on the other hand, depends on theefficiency of the column internals. This traditional method of determining the towerdiameter and height is called the size-based approach.

Another method used in distillation column design is the rate-based approach. However,most commercial columns are still designed using the traditional equilibriumstage/efficiency (size-based) approach. Thus, this guideline only covers the size-based

approach.

Included in this guideline is an example of the Data Sheet generally used in the industryand a Calculation Spreadsheet for the engineering design.

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INTRODUCTION

General Design Consideration

A tower design is normally divided into two main steps, a process design followed by amechanical design. The purpose of the process design is to calculate the number ofrequired theoretical stages, column diameter and tower height. On the other hand, themechanical design focuses on the tower internals and heat exchanger arrangements.

Many factors have to be considered in designing a distillation column such as the safetyand environmental requirements, column performance, economics of the design andother parameters which may constrain the work.

The first step in distillation column design is to determine the separation sequences,which depends on the relative volatility and concentration of each component in the feed.A few design rules have been outlined by King as follows:

1) Direct sequences which remove the components one by one in the distillate aregenerally favored.

2) Sequences which result in a more equal-molar division of the feed betweendistillate and bottoms products should be favored.

3) Separations where the relative volatility of two adjacent components is close tounity should be performed in the absence of other components; ie, reserve such aseparation until the last column in the sequence.

4) Separations involving high specified recovery fractions should be reserved until lastin the sequence.

Once the separation sequence is decided, engineering calculations follow to determinethe number of theoretical stages, operating parameters and tower dimensions. In general,the steps included in distillation calculations are summarized into the following:

1) Performing a material balance for the column2) Determining the tower operating pressure (and/or temperature)3) Calculating the minimum number of theoretical stages using the Fenske equation

4) Calculating the minimum reflux rate using the Underwood equations5) Determining the operating reflux rate and number of theoretical stages6) Selection of column internals (tray or packings)7) Calculating the tower diameter and height

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The theoretical explanation and sample calculations of each step above are discussed indetail in later sections.

Some general design rules (from Cheresources.com) that should be considered are asfollows:

1) Distillation is usually the most economical method of separating liquids.2) For Ideal mixtures (low pressure, medium temperature, and non-polar), relative

volatility is the ratio of vapor pressures i.e. = P2/P13) Tower operating pressure is determined most often by the temperature of the

available cooling medium in the condenser or by the maximum allowable reboilertemperature.

4) Tower SequencingA. Easiest separation first least trays and refluxB. When neither relative volatility nor feed concentrations vary widely, remove

components one by one as overhead products.C. When the adjacent ordered components in the feed vary widely in relative

volatility, sequence the splits in order of decreasing volatility.D. When the concentration in the feed varies widely but the relative volatilities do

not, remove the components in the order of decreasing concentration in thefeed.

5) Economically optimum reflux ratio is about 120% to 150% of the minimum refluxratio.6) The economically optimum number of stages is about 200% of theminimum value.

7) A safety factor of at least 10% above the number of stages by the best method isadvisable.

8) A safety factor of at least 25% about the reflux should be utilized for the refluxpumps.

9) Reflux drums are almost always horizontally mounted and designed for a 5 minholdup at half of the drum's capacity.

10) For towers that are at least 3 ft (0.9 m) in diameter, 4 ft (1.2 m) should be added tothe top for vapor release and 6 ft (1.8 m) should be added to the bottom to accountfor the liquid level and reboiler return.

11) Limit tower heights to 175 ft (53 m) due to wind load and foundation

considerations.12) The Length/Diameter ratio of a tower should be no more than 30 and preferably

below 20.13) A rough estimate of reboiler duty as a function of tower diameter is given by:

Q = 0.5 D2 for pressure distillation

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Q = 0.3 D2

for atmospheric distillationQ = 0.15 D2 for vacuum distillation

where Q is in Million Btu/hr and D is tower diameter in feet.

The Selection of Column Internals

The selection of column internals has a big impact on the column performance and themaintenance cost of a distillation tower. There are several choices of column internalsand the two major categories are trays and packing. The choice of which to utilize

depends on the pressure, fouling potential, liquid to vapor density ratio, liquid loading, andmost importantly the life cycle cost.

Trays can be divided into many categories, such as baffle trays, dual flow trays,conventional trays, high capacity trays, multiple downcomer trays and system limit trays.According to some rules of thumb, trays should be selected if:

1) the compounds contain solids or foulants2) there are many internal transitions3) liquid loads are high4) there is a lack of experience in the service

5) vessel wall needs periodic inspection6) there are multiple liquid phases

On the other hand, packing divisions include grid packing, random packing, conventionalstructured packing, and high capacity structured packing. The rules of thumb for selectingpacking are:

1) the compounds are temperature sensitive2) pressure drop is important (vacuum service)3) liquid loads are low4) towers are small in diameter

5) highly corrosive service (use plastic or carbon)6) the system is foaming7) the ratio of tower diameter to random packing is greater than 10

Some design guidelines should be considered when designing a tray tower, such asfollows:

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1) Tray spacing should be from 18 to 24 inches, with accessibility in mind (Generally,for a tower diameter of 4 feet and above, the most common tray spacing is 24inches to allow easy access for maintenance. However, for a tower diameter below4 feet, a tray spacing of 18 inches is adequate as the column wall can be reachedfrom the manway.)

2) Peak tray efficiencies usually occur at linear vapor velocities of 2 ft/s (0.6 m/s) atmoderate pressures, or 6 ft/s (1.8 m/s) under vacuum conditions.

3) A typical pressure drop per tray is 0.1 psi (0.007 bar)4) Tray efficiencies for aqueous solutions are usually in the range of 60-90% while

gas absorption and stripping typically have efficiencies closer to 10-20%5) Sieve tray holes are 0.25 to 0.50 in. diameter with the total hole area being about

10% of the total active tray area. Maximum efficiency is 0.5 in and 8%.6) Valve trays typically have 1.5 in. diameter holes each with a lifting cap. 12-14

caps/square foot of tray is a good benchmark.7) The most common weir heights are 2 and 3 in and the weir length is typically 75%

of the tray diameter.

The packed tower design concepts are listed below:

1) Packed towers almost always have lower pressure drop compared to tray towers.2) Packing is often retrofitted into existing tray towers to increase capacity or

separation.3) For gas flow rates of 500 ft3 /min (14.2 m3/min), use 1 in (2.5 cm) packing, for gas

flows of 2000 ft3/min (56.6 m3/min) or more, use 2 in (5 cm) packing.4) Ratio of tower diameter to packing diameter should usually be at least 155) Due to the possibility of deformation, plastic packing should be limited to an

unsupported depth of 10-15 ft (3-4 m) while metal packing can withstand 20-25 ft(6-7.6 m).

6) Liquid distributor should be placed every 5-10 tower diameters (along the length)for pall rings and every 20 ft (6.5 m) for other types of random packing.

7) For redistribution, there should be 8-12 streams per sq. foot of tower area fortowers larger than three feet in diameter. They should be even more numerous insmaller towers.

8) Packed columns should operate near 70% flooding.9) Height Equivalent to Theoretical Stage (HETS) for vapor-liquid contacting is 1.3-

1.8 ft (0.4-0.56 m) for 1 in pall rings and 2.5-3.0 ft (0.76-0.90 m) for 2 in pall rings.10) Design pressure drops should be as follows:

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Table 1: Pressure drop in difference services

Service Pressure drop (in water/ftpacking)

Absorbers and RegeneratorsNon-Foaming SystemsModerate Foaming Systems

0.25 - 0.400.15 - 0.25

Fume ScrubbersWater AbsorbentChemical Absorbent

0.40 - 0.600.25 - 0.40

Atmospheric or Pressure

Distillation

0.40 - 0.80

Vacuum Distillation 0.15 - 0.40

Maximum for Any System 1.0

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DEFINITIONS

Bottoms The stream of liquid product collected from the reboiler at the bottom of adistillation tower.

Bubble point The temperature at constant pressure (or the pressure at constanttemperature) at which the first vapor bubble forms when a liquid is heated (ordecompressed).

Dew point The temperature at constant pressure (or the pressure at constanttemperature) at which the first liquid droplet forms when a gas (vapor) is cooled (orcompressed).

Distillate The vapor from the top of a distillation column is usually condensed by a totalor partial condenser. Part of the condensed fluid is recycled into the column (reflux) whilethe remaining fluid collected for further separation or as final product is known as distillateor overhead product.

Equation of state A relation between the pressure, volume and temperature of asystem, from which other thermodynamic properties may be derived. The relationemploys any number of constants specific to the system. For example, for a purecomponent, the constants may be generalized functions of critical temperature, criticalpressure and acentric factor, while for a mixture, mixing rules (which may be dependenton composition or density), are also used.

Heavy key The heavier (less volatile) of the two key components. Heavy key iscollected at the bottoms. All non-key components heavier than the heavy key are knownas the heavy components.

Key component A distillation column is assigned with two key components. The keycomponents in the feed are the main components to be separated in that column. Thevolatility of the two key components must be in adjacent order when the volatilities of allthe components in the feed are arranged in either ascending or descending order.

K-value Vapor-liquid equilibrium constant or distribution coefficient. It is used in non-ideal (hydrocarbon) systems. Some convenient K-value charts are available2, 3.

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Light key The lighter (more volatile) of the two key components. Light key is collectedat the distillate. All non-key components lighter than the light key are known as the lightcomponents.Reflux ratio The ratio of the reflux stream to the distillate. The operating reflux ratiocould affect the number of theoretical stages and the duties of reboiler and condenser.

Relative volatility Relative volatility is defined as the ratio of the concentration of onecomponent in the vapor over the concentration of that component in the liquid divided bythe ratio of the concentration of a second component in the vapor over the concentrationof that second component in the liquid. For an ideal system, relative volatility is the ratio ofvapor pressures i.e. = P2/P1

Vapor pressure The pressure exerted by the vapor phase that is in equilibrium with theliquid phase in a closed system. For moderate temperature ranges, the vapor pressure ata given temperature can be estimated using the Antoine equation.

NOMENCLATURES

b bottoms product flow rate, ft3/min (m3/min)D tower diameter, ft (m)d distillate flow rate, ft3/min (m3/min)F feed flow rate, ft3/min (m3/min)fi fugacity of component iH tower height, ft (m)K vapor-liquid equilibrium constantN number of theoretical stagesNm minimum number of theoretical stagesP total system pressure, psi (bar)P* vapor pressure, psi (bar)Q reboiler duty, Btu/hr (W)R reflux ratioRm minimum reflux ratio

T temperature, 0F (K)x mole fraction in the liquid phasexd mole fraction in the distillatexw mole fraction in the bottomsy mole fraction in the vapor phase

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Greek letters relative volatility activity coefficient vapor phase fugacity coefficient

SuperscriptsL liquid phaseV vapor phase

SubscriptsHHK heavy componentHK heavy keyi component i

j componentjLK light keyLLK light component

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THEORY

A) Vapor-Liquid Equilibrium (VLE)

The design or simulation of a distillation column requires quantitative estimates of thephase equilibrium properties of the mixture. An engineer must be familiar with sources ofequilibrium data, methods for predicting such data when measured values are notavailable, and methods for extrapolating available data.

In a two-phase system in which a vapor mixture is in equilibrium with a liquid mixture, thevariables of interest are the temperature, pressure and the composition of both phases.The condition for thermodynamic equilibrium is given by:

fiV = fi

L Eq (1)

for every component i

The purpose of a VLE model is to relate these fugacities to temperature, pressure andmixture compositions.

The most common methods of representation of VLE data fall into three categories:I) Ideal behavior in both phasesII) Liquid phase non-idealities

III) High-pressure systems

I) Ideal behavior in both phases

Raoults Law can be applied in an ideal system. It is the simplest model, as followed:

fiV = yiP Eq (2)

fiL =xiP*i Eq (3)

where yi= mole fraction of iin vapor phasexi= mole fraction of iin liquid phase

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The only data required for this model are the vapor pressures of the pure components,which can be estimated using an extended Antoine equation.

ln P*i=A +CT

B

++ DT+ Eln T Eq (4)

where Tis the absolute temperature andA, B, C, D, Eare fitted constants.

The Raoults Law model is simple and useful for preliminary feasibility studies. However,for detailed design, this model is only reliable if:

a) the system only contains non-polar compounds, preferably of the same

homologous seriesb) the pressure is low (< 5 bar)c) the relative volatility between the key components is reasonably high (= 3)

II) Liquid phase non-idealities

The liquid phase may show significant deviations from ideality when the componentsbelong to different homologous series, or if it contains polar components. Thesedeviations are expressed in the form of:

fiV

= yiP (the same as before)

fiL = ixiP

*i Eq (5)

where i is the activity coefficient of component i.

Activity coefficients may be greater or less than 1.0. When it is greater, the system isshowing positive deviations from Raoults Law and vice versa. Activity coefficients can bemodeled in two ways, as shown in the following.

Local composition equations

If experimental VLE data are available, they may be fitted by regression to the parametersof a local composition model. A local composition model expresses the activitycoefficient of each component in the liquid phase as a function of composition andtemperature. Local composition models are usually used for moderate or strongly non-

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ideal systems and nearly ideal mixtures with low relative volatility where high accuracy isessential.

Examples of local composition models include Margules, Van Laar, Redlich-Kister,Wilson, NRTL and UNIQUAC. The Wilson and UNIQUAC equations are recommendedfor most applications. However, the Wilson equation is algebraically not capable ofrepresenting a system showing liquid-liquid immiscibility, unless it gives a reasonableapproximation to the average liquid composition on a stage (if the VLE data on either sideof the immiscible region is correlated accurately). UNIQUAC equation is recommendedfor rigorous simulation (if the distillation algorithm can cope with two liquid phases).

Activity coefficient estimation methods

If experimental VLE data are not available, activity coefficients can be estimatedapproximately from the properties of the pure components, such as the solubilityparameter, molecular structure (group contributions), infinite dilution activity coefficients,heat of mixing data and liquid-liquid equilibrium data. The most reliable methods are:

i) solubility parameter methods

These methods are based on Regular Solution Theory and employ theHildebrand expression for activity coefficient as a function of composition andtemperature, using a solubility parameter for each component. This parameteris calculated from molar liquid volume and latent heat of vaporization. Thesemethods are only suitable for non-polar mixtures and are popular in thehydrocarbon industry.

ii) group contribution methods

These methods regard each molecule as an aggregate of functional groups,and evaluate each activity coefficient as a sum of interactions between thegroups. Parameters for each group and for interactions between groups canbe easily found in the literature. These methods (e.g. UNIFAC and ASOG) can

be used for estimation of all organic systems, including strongly non-ideal polarmixtures.

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III) High-pressure systems

At high pressure, the vapor phase also shows significant deviations from ideal behavior.The vapor fugacity may be expressed as:

fiV = i

VyiP Eq (6)

where iV is the vapor phase fugacity coefficient of component i.

Fugacity coefficients are calculated by an equation of state. Some of the most popularequations of state for VLE applications are Redlich-Kwong-Soave (RKS), Peng-Robinson(PR) and Benedict-Webb-Rubin (BWR). RKS and PR are simple and are recommendedfor most applications, especially in systems where an aqueous liquid phase co-exists witha hydrocarbon process phase. The aqueous phase is assumed to be pure water, and thesolubility of water in the hydrocarbon is calculated using a derived interaction parameter.

When a rigorous model is used for the vapor phase, the liquid phase should also bemodeled using sophisticated methods, as in the following.

Vapor fugacity coefficient, liquid activity coefficient

The activity coefficient concept is retained but a rigorous expression for liquid phasefugacity is used as shown below:

fiL = ixifi

*L Eq (7)

fi*L = P*ii

*exp

P

P

L

idP

RT

v

*Eq (8)

where fi*L = fugacity of pure liquid iat Tand Pi

* = saturated vapor fugacity coefficient of pure ivi

L = molar liquid volume of pure i

R = gas constantT = absolute temperature

The exponential term is called the Poynting correction. This model is powerful as it takesinto account experimental data fitted to a semi-empirical equation within a soundthermodynamic framework.

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This design guideline are believed to be as accurate as possible, but are very general not for specific design cases. Theywere designed for engineers to do preliminary designs and process specification sheets. The final design must always beguaranteed for the service selected by the manufacturing vendor, but these guidelines will greatly reduce the amount of upfront engineering hours that are required to develop the final design. The guidelines are a training tool for young engineers

Vapor and liquid fugacity coefficients

In non-polar systems, the liquid phase fugacity may also be represented in terms of afugacity coefficient:

fiL = i

L xiP Eq (9)

iL and i

V can be calculated using the same equation of state. This model is veryconvenient with little input information required, but cannot ultimately match the accuracyof a model based on experimental data. However, equations of state are widely andsuccessfully used in the hydrocarbon industry provided that the relative volatility is high.

The accuracy of this method can sometimes be improved by fitting binary interactionparameters into experimental VLE data. A further development is to fit pure componentvapor pressure data to the basic parameters of the equation of state, and also fit binaryparameters to experimental VLE data. This approach has been applied even to polarsystems at high pressure.

B) Phase Distribution

When a system is in thermodynamic equilibrium, the vapor and liquid phases show adistribution volatility between pairs of components. This separation factor has itsequivalent term in other methods of separation that are based on the equilibrium concept.For components iandj, their relative volatility is defined as:

ij= Ki/Kj Eq (10)

Each component iof the mixture has a distinct vapor-liquid equilibrium ratio or K-value.

Ki= yi*/xi Eq (11)

where yi* is a vapor mole fraction, in equilibrium with liquid mole fractionxi.

K-value can be estimated using the thermodynamic relationships as discussed earlier.For example, the K-values for high-pressure systems take the general form of:

Ki= iP*ii

*exp

P

P

L

idP

RT

v

*/i

VP Eq (12)

20 design

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