Andreiadis Design of a Rectification Unit

8/12/2019 Andreiadis Design of a Rectification Unit

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UnitOpe

rationsProject

Politehnica University of BucharestFaculty of Engineering in Foreign LanguagesChemical Engineering Division

Design of a

Rectification Unit

Eugen S. AndreiadisGroup 1244 E

Year 2003 - 2004


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Unit Operations Project Design of a rectification unit

2

Foreword

The objective of this project is the calculation of two rectification towers, oneemploying short-cut methods and the second using rigorous methods.

The project begins with a brief presentation of the method of separation

employed (i.e. rectification), its advantages over distillation and the problemswhich may arise in multi-component rectification.

Following, the first column calculation is done, starting with mass balances andthe determination of the temperature profile, and continuing with the minimumreflux ratio (by the Underwood method) and the total reflux ratio (employingthe Fenske equation). An optimal reflux ratio and the corresponding theoreticalnumber of trays are obtained by the Galliland method, after which all thepreviously computed values are verified against a Hysys simulation, thusfinalizing the first column calculation.

The second column is assumed to be fed with a binary mixture, simplifying inthis way the computations. This more rigorous analysis also includes ahydrodynamic calculus and the determination of the mass transfer coefficientsand the real number of trays. Knowing this specific details, it is possible to drawa scaled representation of the second column.

The project ends with several Appendixes containing raw data extracted fromHysys concerning the first and also the second column simulations.


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Project Outline

Project Brief 4Signification of Symbols 5

1. Justification 62. First Column Calculation 9

2.1. Mass Balances 9 2.2. Temperature Profile 11 2.3. Minimum Reflux Ratio 14 2.4. Total Reflux Ratio 15 2.5. Optimum Reflux Ratio 16 2.6. Hysys Simulation 17

3. Second Column Calculation 193.1. Theoretical Number of Trays 19

3.2. Hydrodynamic Calculus 22 References 33 Appendixes 34


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Signification of Symbols

F feed stream or feed flowD distillate stream or distillate flowW waste stream or waste flowyi vapour composition of component ixi liquid composition of component iij relative volatility of i to jsi degrees of recoveryLKC light key componentHKC heavy key componentKD,i distribution coefficient of component iPr reduced pressure

Pc critical pressurePv vapour pressureTr reduced temperatureq thermal state of feedLmin minimum reflux ratio (infinite number of trays)Nmin minimum number of theoretical trays (infinite reflux ratio)NT theoretical number of trays (optimum reflux ratio)QV, QL vapour or liquid flow density

dynamic viscosity

surface tensionw vapour or liquid velocity

Di diameter of the columnSr free area of the tray (hole area)S transversal section areaSd downcomer areaSa active areaz number of valveshz weir heightH tray spacingt valve spacingld weir lengthde equivalent diameterdS valve tray diameterd

0

valve hole diameterhS maximum valve heighte liquid entrainmentG1 valve mass

S tray thickness

DV, DL diffusion coefficientskx, ky partial mass transfer coefficientsK total mass transfer coefficientRe Reynolds criterionSc Schmidt criterionSh Sherwood criterion


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1. Justification

Separation of individual substances in a homogeneous liquid mixture or completefractionation of such mixtures into their components is an important step in manyproduction processes. Different separation procedures can be used for this purpose,but distillation is the most important industrial method.

Distillation utilizes a very simple separation principle based on the development ofintimate contact between the homogeneous mixture and a second phase, whichthereby allows mass transfer to occur between phases. The thermodynamic conditionsare chosen so that only the component to be separated enters the second phase. Thephases are subsequently separated; one of them contains the desired substance, and

the other consists of a mixture that is largely free of this substance.

Three steps are always involved in industrial implementation of this separation principle:

Creation of a two-phase system Mass transfer between phases Separation of the phases

A large number of separation techniques utilize this very effective principle ormodifications thereof. Absorption, desorption, evaporation, condensation, and

distillation involve a gaseous and a liquid phase; solvent extraction uses two liquidphases. Separating techniques that utilize a fluid phase and a solid phase areadsorption, crystallization, drying, and leaching. In most of these separations, thenecessary two-phase system is created by adding an auxiliary phase to the mixture;the diluted substances to be separated collect with the auxiliary agent. However, indistillation, the second phase is produced by partial vaporization of the mixture.Hence, the use of an auxiliary substance, which usually requires laborious recovery,can be avoided, and the components to be separated can be recovered as puresubstances. Indeed, distillation requires energy only in the form of heat, which cansubsequently be removed from the system.

The vapour and liquid are brought into intimate contact by countercurrent flow andmass exchange occurs because the two phases are not in thermodynamic equilibrium.The phases produced during rectification are formed by evaporation andcondensation of the initial mixture. The separation of a liquid mixture into its purecomponents can be controlled solely via the heat supply.

The basis of this unit operation is the volatility difference between the properties ofthe liquid and the vapour phase, respectively, on the vapour-liquid equilibrium of thesystem.

Among the main methods used in distillation practice (differential, flash, batch,azeotropic, extractive distillation, a.s.o.), we are interested here solely in rectification.


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Figure 2.Multiple distillation of a binary mixture a-bFlow diagram with (A) condensation or without (B) condensation

Figure 1. Continuous distillation ofa binary mixture a-b

A) Flow diagram showing symbols for totalmolar streams or flow rates (F, L, etc.) and

mole fractions of the more volatile component

a (xF, xD, etc.); B) A y-x diagram showingequilibrium and operating lines

In differential or flash distillation, the vapour leaving the still at any time is inequilibrium with the remaining liquid, and there will normally be only a small increasein the concentration of the more volatile component. The essential merit ofrectification is that it enables the vapour to be obtained to be substantially richer

than the liquid left in still. This is achieved by an arrangement known as fractionatingcolumn which enables successive vaporisation and condensation to be accomplishedin one unit.

Figure 1 and 2 illustrate the differences between

continuous distillation and multiple distillation,i.e. rectification.

The fractionating column consists of a cylindricalstructure divided into sections by a series ofperforated trays which permit the upward flowof vapour. The liquid reflux flows across eachtrays over a weir and downcomer to the traybelow. The vapour rising from the top traypasses to a condenser and then to some form

of reflux divider where part is withdraw as aproduct D and the remainder is returned to thetop tray as reflux. The reflux stream isfrequently passed from the condenser througha reflux drum and then pumped to the column at a rate determined by a suitablecontrol device. The liquid in the base of the column is heated, either by condensingsteam or by a hot oil steam, and the vapour rises through the perforations to thebottom tray.

A more commonly used arrangement consists of an external reboiler. Here theliquid from the still passes into a reboiler where it flows over the tubes and leaves as

the bottom product; the more volatile material returns as vapour to the still.


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Vapour of composition ywenters the bottom tray (say n) where it ispartially condensed and then revaporised to give vapour ofcomposition yn. This operation of partial condensation and partialvaporisation of the reflux liquid id repeated on each tray.

The feed stream is introduced on some intermediate tray wherethe liquid has approximately the same composition as the feed.The part of the column above the feed point is known as therectifying section, while the lower part is known as the strippingsection.

Figure 3. Flow of vapour and liquidthrough a rectification column

Multi-component mixture rectification is a more frequently operation in industrialchemistry and refineries compared with a binary one, and this project implies the

existence of such a mixture, having 3 components. One of the main problems such anapproach exhibits is the optimal selection and sub-sequencing of separationoperations and adequate equipment in order to meet the specified processrequirements.

For example, let us consider such a mixture of 3 components A, B and C (given inorder of decreasing volatility) which should be separated until a given purity bycontinuous rectification. Three schemes can be imagined, as follows from Figure 4.

Scheme 1 Scheme 2 Scheme 3

Figure 4.Sequencing of rectification towers

Scheme 1 uses the first column for obtaining A like distillate. This fraction still containssome quantities of B and C (only traces, and if the relative volatility of A is big enough,these quantities can be neglected). The waste fraction (a binary mixture of B and C,

contaminated with small quantities of A) is separated in the second column.

A+B+C

A+B+C

(A+)B(+C)A+B+C

1

2

1

2

1

2

A(+B)

(A+)B+C

(A+)B

(B+)C

A+B(+C)

(B+)C

A(+B)

B(+C)

A(+B)

(B+)C


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Scheme 2 uses the first column for obtaining C like waste (in the bottom). Thedistillate, also containing a small quantity of C, is separated further in the secondcolumn in order to obtain two fractions corresponding of A and B. According to therelative volatility, the component C cannot be found in the distillate of the secondcolumn.

The third scheme is much more complicated, because of the side streams in bothcolumns, but is also more economic.

In practice, choosing one of the schemes is a matter of costs. The first scheme is moreeconomic than the first because it boils only one component, A (assuming A is liquid).On the other hand, when the third component C is corrosive or toxic, it is better toeliminate it first. Anyway, one column can reasonably isolate only one component, sofor ncomponents we have n-1columns.

The sequencing scheme imposed for our project is the second, so we are to separatecomponent C in the first column and then isolate A and B in the second one.

2. First Column Calculation

The feed introduced in the first column contains three components. An importantapproximation that has to be made concerns the so called key components. Thevolatile components are called light and the less volatile ones are called heavy. Thelight key component (LKC) will be considered to be that light component which is

found as an important fraction of the waste flow (all the lighter ones could beneglected). If all light components are significant fractions of the waste, the lightestone will be called the light key component. The heavy key component (HKC) is thatheavy component which could be found as an important amount in the distillate (orthe heaviest one, if all key components are present in the distillate).

LKC: component 2 (propane)HKC: component 3 (butane)external: component 1 (ethane)

In our case (scheme 2), the LKC is considered to be component 2, while the HKC is

considered to be component 3. Assuming that the keys were chosen properly (whichwill be verified later using the Shiras, Hanson and Gibson equation), the calculusevolves, with some modifications, like for a binary mixture composed of these keycomponents.

2.1. Mass Balances

A simple representation of the flows and the composition of the flows in the case ofthe first column is given below in Figure 5.

The separation degrees are as follows


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2F

2Dv

xF

xDs

(1)

3F

3Wh

xF

xWs

(2)

and the key component approximation actually implies

0x 1W (3)

The mass balances are

WDF (4)

1W1D1F xWxDxF (5)

2W2D2F xWxDxF (6)

The system above contains only six equations but should

calculate eight unknowns. We also have to add the stoichiometric restrictions

1xDj (7)

1xWj (8)The computations were done first by hand and then were verified in an Excel sheet.The results obtained after solving the system are presented below

Similarly we can compute the mass balance for the second column, taking intoaccount the differences between the equations of the system and the fact that thedistillate from the first column becomes the feed for the second.

1F

1Dv

xF

xDs

(1)

2F

2Wh

xF

xWs

(2)

0x 3D (3)

WDF (4 )

2W2D2F xWxDxF (5)

3W3D3F xWxDxF (6)

F L D

V

L

V

W

Figure 5. First Column

Data for the mass balance: F, xFj, sv, sh sv (%) = 99 sh(%) = 98

Calculus hypothesis: xW1= 0

Flows (kmol/h): Concentrations 1 2 3 Verification:

F= 950 xF= 0.2 0.3 0.5 1

D= 481.65 xD= 0.3945 0.5858 0.0197 1

W= 468.35 xW= 0.0000 0.0061 0.9939 1

Total balance: 0 Partial balances 0 2E-14 0


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1xDj (7 )

1xWj (8 )We get the values

We see that in both cases the mass balances close, although for some columns theremay be obtained instead of zero very small values, because of the truncation errors.

2.2. Temperature Profile

The temperature in the column is needed in order to verify the componentsdistribution. The temperatures of the distillate, the feed and the waste are dependent

upon the type of condenser and reboiler used and also upon the thermal state of feed q.

Top and Bottom Temperatures

The temperature at the top of the column is dependent upon the type of condenseremployed. Thus, for the case of a total condenser, it means that the vapour fractionon the first theoretical stage has a known composition yDi = xDi. The pressure beinggiven, a Dew T calculus with this composition gives the desired temperature.

The temperature at the bottom of the column depends on the type of reboiler used.If the reboiler works at equilibrium (it is considered a theoretical plate), the reboiling

ratio W/V is needed to estimate the temperature (see Figure 5). The real temperatureis between the boiling point of the liquid fraction W and the temperature of thevapour which enters the N+1 stage. In order to simplify the calculus, the dew point ofW is considered to be the searched temperature.

The calculations are done employing a Pascal program which tries severaltemperatures until the needed conditions are fulfilled (verification is done using theRiedel Plank Miller relationship, see below). We found the values

TD= 29.21 C (for the top of the column)

TW= 95.85

C (for the bottom of the column)

Data for the mass balance: F, xFj, sv, sh sv(%) = 95 sh(%) = 97

Calculus hypothesis: xD3= 0

Flows (kmol/h): Concentrations: 1 2 3 Verification:

F= 481.65 xF= 0.3945 0.5858 0.0197 1

D= 188.975 xD= 0.95521 0.044792 0 1

W= 292.675 xW= 0.03246 0.935119 0.03242 1

Total balance: 0 Partial balances: 0 0 0


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Verification of Computed Values

The two values obtained have been verified using the De Priester diagram and alsothe Riedel-Plank-Miller relationship. In the case of the Dew T calculus, the mole

fractions in the vapour phase (yi) are known, we need to verify if indeed the estimatedvalues for Kdiverify the relation

1K

yx

Di

ii (9)

For the case of the Bubble T calculus, the relation to be used is

1Kxy Diii (10)The Riedel-Plank-Miller relationship states that

3

i,ri

2

i,r

i,r

ivi,r T1gT1T

G

Plog (11)

where Pr,vi is the reduced vapour pressure of component I, T r,I is the reducedtemperature and Gi and gi are two constants specific for each component. Therelation allows the estimation of Pvsimilarly to the Antoine equation, only that it offersmuch more confidence at higher pressures because of the presence of reduced(involving critical) parameters.

The computations were done using Mathcad, an example of the results obtained inthe case of n-Butane being presented below

G 1.50014 g 1.81934 temp redusa = temp / temp criticapres redusa = pres / pres critica

Tcr 425.12 Pcr 37.34 Tr302.36

Tcr etan Tcr = 32.3 C = 305.45 K

Pcr = 48.2 atm

propan Tcr = 96.68 C = 369.83 K

Pcr = 41.92 atm

n-butan Tcr = 151.97 C = 425.12 K

Pcr = 37.34 atm

lgG

Tr

1 Tr2 g 1 Tr ( )3

lg 1.135

Parametrii Miller (G si g)

C2H6 1.38881 1.55945 etan

C3H8 1.44933 1.69639 propan

nC4H10 1.50014 1.81934 n-butan

Pvap 10lg

Pcr

Pvap 2.739

KdPvap

13.82

from Lange's Handbook of

Chemistry, 15th Ed, John A. DeanKd 0.198

The results are systematised in Table 1. We see that the temperatures are very well

verified, as it was to be expected since the Pascal program computed them on thebasis of the same equation.


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Table 1. Riedel Plank Miller validation

T (C) Comp K Di xDi Dix1 3.268 0.12072 0.751 0.780029.21

3 0.198 0.0995

1.0023

1 12.213 0.00002 2.988 0.018295.853 0.996 0.9899

1.0081

The results obtained using the De Priester diagram are also given in Table 2. Thevalues obtained are also in good agreement with the expected ones.

Table 2. De Priester validation

T (C) Comp K Di xDi

Dix

1 2.55 0.15472 0.82 0.714429.213 0.26 0.0758

0.9449

1 5.15 0.00002 2.4 0.014695.853 1.0 0.9939

1.0085

Feed Temperature

The temperature of the feed stream is a function of its thermal state q. This parameteris defined as the ratio between the liquid part of the feed and the total feed,

F

Fq L (12)

The calculus is done using the same algorithm as before and the same Pascalprogram, the condition to be fulfilled in this case being

1)q1(Kqx

xi,D

i,F*

i,F (13)

where*

,iFx denotes the composition of the liquid part of the feed. We have obtained

TF= 60.21 C

and the following phase composition

*

1,Fx = 0.0523*

1,Fy = 0.3208*

2,Fx = 0.2347*

2,Fy = 0.3534*

3,Fx = 0.7129*

3,Fy = 0.3251

The average temperature in the column is computed as an arithmetic mean betweenthe top and the bottom temperature, giving

Tm= 62.53 C


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2.3. Minimum Reflux Ratio

The minimum reflux ratio requires an infinite number of theoretical plates in order toseparate the key components as required, so it is a value that cannot be practicallyachieved, and any reflux ratio has to be grater than this quantity.

Relative Volatilities

Employing the same Riedel-Plank-Miller relation, this time for the mean temperaturein the column, we can obtain the distribution coefficients KD,i and from here therelative volatilities (given in respect to the HKC) according to

HKC,D

i,D

iK

K (14)

The results are presented in Table 3.

Table 3. Relative volatilities

T (C) Comp K Di i1 1.579 13.2752 6.412 3.26962.533 0.483 1

Shiras, Hanson and Gibson Validation

The calculus done using the minimal reflux hypothesis allows us to validate theassumptions made regarding the key components: all internal components should bepresent in significant amounts in both effluents and could be considered asdistributed ones. Shiras, Hanson and Gibson have shown that, for a minimal reflux, anapproximate relation is generally valid,

Fx

Dx

1Fx

Dx

1

1

Fx

Dx

HKC,F

HKC,D

LKC

iLKC

LKC,F

LKC,D

LKC

i

i,F

i,D

(16)

where the relative volatilities i are calculated as above. If

Fx

Dx

i,F

i,D

< -0.01 or Fx

Dx

i,F

i,D

> 1.01,

then the component i is probably not distributed, but if

99.0,01.0Fx

Dx

i,F

i,D

,

then i is distributed for sure (is not external). In our case the above formula is appliedonly once in order to verify if the first component is really external. We obtain

Fx

Dx

i,F

i,D

= 5.268 > 1.01

and so indeed component 1 is not distributed.


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Underwood Method

The value of the minimum reflux ratio Lminis only needed to verify the real reflux ratioand thus a very accurate method is not required. We chose the algorithm proposed

by Underwood, which considers constant the relative volatilities i and the ratio L/Vand assures a reasonable value after a moderate effort. The following system has tobe solved.

1LDxD

q1FxF

min

i

i,Di

i

i,Fi

(17-18)

Equation (17) should be written for all components i from the feed and should be

solved with respect to . Each solution obtained (between 1 and LKC ) gives a variantof equation (17), which is also written for all components i from the distillate.

The system was solved employing a Pascal program. The correct value for waslocated using the bisection method. We obtained = 2.037 and thus

Lmin= 1.0012

2.4. Total Reflux Ratio

An infinite reflux ratio implies that there is no distillate, all the top product beingreturned in the column. Since the products distribution changes with the reflux ratio,the calculus at total reflux could give precious information regarding the finalcompositions. In order to find the minimum number of trays corresponding to theinfinite reflux, we use an extended form of Fenske equation, written for keycomponents,

LKC,W

HKC,W

HKC,D

LKC,D

LKC

minxW

xW

xD

xDlog

log

11N (19)

where Nmin denotes the minimum number of plates. If the condenser and/or thereboiler work at equilibrium, they are included in Nmin. The average relative density iscomputed as a geometrical mean

WD T

LKC

T

LKCLKC (20)

We find, employing the same Riedel-Plank-Miller relation, DTLKC = 3.793,WT

LKC = 3.000

and thus LKC = 3.373, following that

Nmin= 6

Equation (19) can also be used to calculate the distribution of other components thatthe key ones in the case of total reflux.


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HKC,W

HKC,D1N

LKC

i,W

i,D

xW

xD

xW

xDmin

(21)

In this case, only the external 1 components distribution could be calculated. This willgive a correction of the mass balance, allowing for more accurate compositions and

flows. For this, the previous assumption xW,1 = 0 will be replaced by the valueobtained according to (21). The calculations are done using the following Excel sheet.The value denoted with alfa is the right side of equation (21). Of course, it is stillpossible to obtain non-zero values for some mass balance verifications.

2.5 Optimum Reflux Ratio

This parameter is to be estimated using the method proposed by Gilliland, which

gives confident results for different thermal states q and a large range of relativevolatilities. The following correlation has been made,

1X002743.0X591422.0545827.0Y (22)

where2N

NNY

T

minT

and1L

LLX min

.

For several values of the reflux ratio L (L >Lmin), X is calculated with the above formulaand Y from equation (22). From this, agraphical dependence of the type NT= f(L)

can be obtained. The optimal reflux ratio isconsidered to be that value where thecurve tends to remain constant (obtainedby drawing two tangents to the curve andlocating their interception point). Thecomputations were done using an Excelsheet and the results obtained are plottedin Figure 6. A part of the data streamemployed is given in the sheet aside.

It follows that Lopt= 3 andNTT = 9.

L X Y N

1.02 0.009307 0.835049 46.50

1.50 0.199520 0.441574 12.33

2.00 0.332933 0.357162 10.44

2.50 0.428229 0.298969 9.413.00 0.499700 0.255783 8.75

3.50 0.555289 0.222357 8.29

4.00 0.599760 0.195689 7.95

4.50 0.636145 0.173908 7.68

5.00 0.666467 0.155780 7.48

5.50 0.692123 0.140453 7.31

6.00 0.714114 0.127325 7.17

Data for the mass balance: F, xFj, sv, sh sv (%) = 99 sh(%) = 98

Calculus hypothesis: D*xD1=alfa*W*xW1 a1 = 3.373 Nmin= 6

alfa = 101.37

Flows (kmol/h): Concentrations 1 2 3 Verification:F= 950 xF= 0.2 0.3 0.5 1

D= 479.79 xD= 0.3921 0.58806 0.0198 1

W= 470.21 xW= 0.0039 0.00606 0.98999 1

Total balance: 0 Partial balance -7E-15 2E-14 0


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5.80

6.80

7.80

8.80

9.80

0.00 5.00 10.00 15.00 20.00 25.00 30.00 35.00 40.00 45.00 50.00 55.00

L

N

t

Figure 6. The Gilliland method for obtainingthe optimum reflux ratio

2.6 Hysys Simulation

Following the determination of the optimum reflux ratio, we chose to verify all theabove calculations using a recent version (3.0.1) of the powerful software packageHysys, the most famous process simulator. The chosen method was the so-calledshort-cut distillation, providing initial estimates for a much rigorous later calculus.

This method is based on the Fenske-Underwood equations, and it gives Nminand Lmintogether with the number of ideal trays, the optimal feed location and other variousproperties for the components in the mixture.

Procedure

The procedure goes as follows: firstly, from the list of available components arechosen the ones to be separated. For each a set of physical and thermodynamicalproperties can be consulted. For reference, the properties found for our componentsare listed in the Appendix. The next step requires the selection of one Fluid Package

to be used, and we have selected the PRSV package (based on the Peng-Robinsonequation of state).


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Now the access to the simulation manager is granted. This visual environment allowsus to select and connect the equipments that are going to be used. Thus, we haveintroduced the two columns and joined them through a valve (since there exists apressure difference between the two). We were required to name each flow rate andto specify the pressure at the top and the bottom of the column (the same in the case

of the first column and in the first approximation for the second one). We could haveintroduced the previously determined temperatures at the two ends, but we havechosen not to do so and let the simulator compute them. The reflux ratio L as well asthe liquid rate can also be specified in advance. Instead, we have introduced theimposed recovery rate svand sh. More information are needed to fully characterise thefeed stream F, for example the composition of the stream xFi, the pressure, thethermal state q and of course the molar flow rate. At this moment the simulator cancalculate all the other characteristics of the streams and the column that we areinterested in.

Data inputs and outputs

F = 950 kmole/hxF1= 0.2xF2= 0.3P = 1400 kPaq = 0.45 (q = 0.55)

Having introduced the above inputs, we were able to try next several values for thenumber of plates until the value for the reflux ratio approached the Lopt computedwith the Galliland method. However, for the NTT = 9 determined at paragraph 2.5.,

the column didnt converge. The convergence was assured only for values of NTTgreater than 11. Moreover, the Lopt= 3 was reached only at much higher number oftrays, as can be seen from the Table 4.

Table 4. First column

Note that since the reboiler works at equilibrium, it isconsidered a theoretical tray also.

The location of the feed point was chosen as to correspond tothe same composition of the vapour-liquid system inside thecolumn as in the feed stream. For this, the compositions of the

liquid key components on each tray were extracted and thendivided. Knowing the thermal state of the feed

i,F

L

i,F

x

xq (23)

the compositions of the liquid key components in the feed can be obtained and theirratio calculated. This ratio is compared then with the ratio on each tray and theclosest resemblance belongs to the feed tray.

After several trial-and-error attempts, we have chosen a NTT = 17 corresponding to a

reflux ratio L = 2.84. All the important values obtained after running the simulation inthis conditions are presented in the Appendix.

NT Nfeed L11 5 21.412 5 8.0813 6 5.4314 6 4.1815 6 3.5216 6 3.12

17 6 2.8418 7 2.6219 7 2.4620 7 2.34


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The same was applied for the second column, as we tried to find the ideal number oftrays and the location of the feed point that would minimize the reflux ratio, in thesame time keeping a smooth temperature profile in the column. The optimum resultsobtained by such trial-and-error method for the both columns are resumed in the

table below.

Table 5. Hysys simulation. Trial-and-error method

Column NT Nfeed L1 17 6 2.842 9 4 1.43

3. Second Column Calculation

If the first column was fed with a ternary mixture, this one is considered to be a binarycolumn, in order to simplify the calculations. For this, the component 3 wasdistributed so that the compositions of the first two would total 1. Also, a new columnwas simulated in Hysys according to the conditions above. Employing the simulator,various physical and thermo-dynamical properties for the selected components (1 and2) were found in order to be used later in the hydrodynamic calculus.

3.1. Theoretical Number Of Trays

The column calculation was done with the help of an Excel sheet, in which severalfields needed to be introduced in order to obtain the mass balance, Lmin, Nmin andNTT, as detailed in the case of the first column. The results are presented below.

After obtaining the mass balance, the liquid-vapour diagram is to be plotted. Forseveral values of temperature T in the range of boiling points of the two components,a special form of the Antoine equation (having 6 parameters for more accuratepredictions at higher pressures) was employed to give the vapour pressures for thetwo components. (Note: for the calculation of the ternary column, we have used

instead the Riedel-Plank-Miller relation.) The equation is as follows

Components Molar Masses Antoine Constants

A B C D E F

C2H6 etan 30 kg / kmol 44.0103 -2568.82 0 -4.97635 1.46E-05 2

C3H8 propan 44 kg / kmol 52.3785 -3490.55 0 -6.10875 1.12E-05 2

Technological Imputs Mass balance Boiling Points

x1 x2

F = 481.67 kmol / h F = 481.67 0.402441 0.597559 1 -39.23

xF = 0.402441 kmol 1/ kmol am D = 192.7791 0.955253 0.0447466 1 18.1xD = 0.955253 kmol 1/ kmol am W = 288.8909 0.033545 0.9664548 1

xW = 0.033545 kmol 1/ kmol am 0 -4.09E-14 0

p = 800 kPa

q = 0.857


20/51


20

f

V TeTlndcT

baPln

(24)

where the pressure is obtained in kPa, the temperature is introduced in K and theparameters a to f have to be obtained from tables (or the simulator).

Knowing the vapour pressures, it is easy to calculate the corresponding molarfractions x and y with the formulas

2,V1,V

2,V

1PP

Ppx

(25)

p

Pxy 111 (26)

where p is the total pressure, and PV,i are the vapour pressures of components icalculated with the Antoine equation. Having computed these, it is easy now to get

the values of the relative volatility with

2,V

1,V

2,1P

P (27)

and to mediate them, obtaining finally an average relative viscosity. Also, theequilibrium curve can be plotted as y = f(x) (see below).

Next the minimum reflux ratio is computed as

F

*

F

*

FDmin

xy

yxL

(28)

where yF*represents the equilibrium concentration corresponding to xF.

The optimal reflux is computed with the Woinaroschy equation

Liquid-Vapour Equilibrium

n T pV1 pV2 x y alfa12

0 -39.23 799.6177 115.0705 1.000558 1.00008 6.9489361 -36.3635 876.7525 129.9149 0.89723 0.983311 6.748666

2 -33.497 959.2057 146.1995 0.804177 0.964213 6.560936

3 -30.6305 1047.196 164.0144 0.720107 0.942617 6.384781

4 -27.764 1140.946 183.4517 0.643918 0.918345 6.219327

5 -24.8975 1240.683 204.6055 0.574662 0.891217 6.063783

6 -22.031 1346.638 227.5715 0.511523 0.861046 5.917429

7 -19.1645 1459.046 252.4472 0.453798 0.827641 5.779609

8 -16.298 1578.149 279.3319 0.400879 0.790808 5.649728

9 -13.4315 1704.192 308.3261 0.352236 0.750347 5.527239

10 -10.565 1837.427 339.5322 0.30741 0.706054 5.411645

11 -7.6985 1978.114 373.0538 0.266 0.657723 5.302489

12 -4.832 2126.515 408.9961 0.227656 0.605143 5.199353

13 -1.9655 2282.905 447.4658 0.192071 0.548099 5.101854

14 0.901 2447.562 488.5708 0.158974 0.486374 5.009637

15 3.7675 2620.775 532.4204 0.128129 0.419748 4.92238

16 6.634 2802.841 579.1256 0.099327 0.347996 4.839781

17 9.5005 2994.066 628.7986 0.072381 0.270893 4.76156618 12.367 3194.766 681.5532 0.04713 0.18821 4.687479

19 15.2335 3405.269 737.5046 0.023426 0.099715 4.617285

20 18.1 3625.912 796.7699 0.001142 0.005175 4.550764

5.533556

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1


21/51


21

min

min

min

opt

L

N192.01

L

L (29)

while the minimum (Nmin) and the theoretical number of trays (NTT) are determined bythe graphical McCabe-Thiele method. The equation for the rectification line is

DD xx1L

Lxy

(30)

and for the stripping line

WW xx1L

FLxy

(31)

while the Fenske equation is

y1y

x

(32)

Minimum reflux ratio

yf* = 0.788436 kmol 1 / kmol am

Lmin = 0.432174

Minimum number of trays

n x y

0 0.955253 0.955253

1 0.794151 0.794151

2 0.41079 0.41079

3 0.111895 0.1118954 0.022262 0.022262

Nmin = 4

Optimum reflux

Lopt = 0.684615

QL = 131.9794 kmol / h

QV = 324.7585 kmol / h

QL ' = 544.7706 kmol / h

QV ' = 255.8797 kmol / h

xq = 0.351162 kmol 1 / kmol am

Theoretical number of trays

n x y

cond 0.955253 0.955253

1 0.794151 0.955253

2 0.593316 0.889783

3 0.432244 0.808165

alim 4 0.342821 0.742706

5 0.288771 0.691997

6 0.19771 0.576924

7 0.100884 0.383053

8 0.03739 0.176911

9 0.007808 0.04173

NTT = 9

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1


22/51


22

After knowing the NTT and the location of the feed point, it is possible to run thesimulator and obtain various physical and thermo-dynamical data for each tray (seeTable 6). The values will be further used in the hydrodynamic calculus.

Table 6. Physical properties for the second column

3.2. Hydrodynamic Calculus

General Considerations

The process of distillation can carried out in many types of columns, the mostcommon being the plate and the packed columns. Our project was imposed to have arectification tower with plate trays, more specifically valve trays. Here the separation iscarried out in a stage-wise manner, whereas with packed columns the process ofmass transfer is continuous.

The number of theoretical stages required to effect the desired separation and thecorresponding rates for the liquid and vapour phases was determined by theprocedures described earlier, but to translate these quantities into an actual design,the following factors have to be considered

the type of tray (in our case valve tray) the vapour velocity (the major factor in determining column diameter) the plate spacing (the major factor fixing the height of the column when the

real number of stages is known)

The main requirement for a tray is that it should provide intimate mixing between theliquid and the vapour streams, as well as being suitable of handling the desired rates

without excessive entrainment or flooding, being stable in operation and being easyto erect and maintain.

Stage Surface Tension Mole Weight (Vap) Density (Vap) Viscosity (Vap)

Condenser 8.480 dyne/cm 30.19 14.38 kg/m3 7.732e-003 cP

1__Main TS 8.769 dyne/cm 30.7 14.24 kg/m3 7.855e-003 cP








Reboiler 7.895 dyne/cm 42.66 16.81 kg/m3 8.390e-003 cP

Stage Surface Tension Mole Weight (Lt Liq) Density (Lt Liq) Viscosity (Lt Liq)

Condenser 8.480 dyne/cm 30.7 480.6 kg/m3 9.783e-002 cP

1__Main TS 8.769 dyne/cm 32.78 492.6 kg/m3 0.1044 cP2__Main TS 9.001 dyne/cm 35.78 504.9 kg/m3 0.1110 cP

3__Main TS 8.978 dyne/cm 38.15 510.3 kg/m3 0.1143 cP






Reboiler 7.895 dyne/cm 43.63 504.9 kg/m3 0.1081 cP


23/51


23

The most important types of trays are as follows

bubble cap trays sieve trays

valve trays

The latter type, which may be regarded as intermediate between the bubble cap andthe sieve cap, offers advantages over both. The feature of the tray is that liftable capsact as variable orifices. A diagrammatic representation of a valve is given in Figure 7.

The valves (1) are metal disks or metalstripes of up to 38 mm in diameter whichare raised above the openings (3) in thetray deck (2) as vapour passes through thetray. The caps are restrained by legs orspiders which limit the vertical movementup or down. Some types are capable offorming a total liquid seal when the vapourflow is insufficient to lift the cap.

A cutaway section of a plate column is given in Figure 8, while a schematicrepresentation of the most important parameters characterising a tray is representedin Figure 9.

Data Inputs

The computations below will be made separately (when it is necessary) for therectifying and the stripping region of the column, at the values corresponding to themedian trays. Some important data used in the calculus are gathered in the Table 7for quick reference.

Table 7. Data inputs for the hydrodynamic calculus

Parameter Symbol Rectification Stripping

Median tray - 2 7

Vapour flow QV 324.76 kmole/h729.00 m3/h 255.88 kmole/h629.42 m3/hLiquid flow QL 131.98 kmole/h

9.33 m3/h544.77 kmole/h44.26 m3/h

Temperature t -24.2 C 1.2 C

Vapour mole weight MV 31.88 kg/kmole 37.66 kg/kmole

Liquid mole weight ML 35.68 kg/kmole 41.46 kg/kmole

Vapour densityV 14.18 kg/m

3 15.31 kg/m3

Liquid densityL 504.6 kg/m

3 510.3 kg/m3

Vapour dynamic viscosityV 8.05 E-3 cP 0.1108 cP

Liquid dynamic viscosity L 8.392 E-3 cP 0.1128 cP

1

5

2

3

Figure 7.A valve


24/51


24

Figure 8.Section through a plate columna) Downcomer; b) Tray support; c) Sieve trays; d)Manway; e) Outlet weir; f) Inlet weir; g) Side wallof downcomer; h) Liquid seal

Figure 9.Schematic of a plate column showing important parameters

A) Vertical section: a) Bubble-cap plate; b) Valve plate; c) Sieve platedcap

= cap diameter;

dv= valve diameter; d

h= hole diameter; h

cl= height of skirt clearance of downcomer; h

w= weir height; H =

plate (tray) spacing; VL= volumetric flow rate of liquid

B) Cross section:d) Small holes with narrow spacing; e) Large holes with wide spacing;

Ad

= cross-sectional area of downcomer; Aac

= active cross-sectional area; Dc

= column diameter; lcl

= skirt

clearance length of downcomer; l = length of liquid path; l = weir length


25/51


25

We have chosen the following characteristics for our type of tray:

Table 8.Chosen parameters

Parameter Symbol ValueValve tray diameter dS 75 mmValve hole diameter d0 65 mmValve weight G1 80 gMaximum valve height hS 12 mmValve material - CuValve material density

Cu 8900 kg/m3

Tray spacing H 0.3 mTray thickness

S 5 mm

Hole area Sr 0.15 S

Vapour velocity

The vapour velocity in a valve column can be determined from a graphicaldependence Y = f(X) corresponding to the optimum operating conditions. Here

125.0

V

L

25.0

V

L

Q

QX

(33)

16.0

V

L

L

V

2

re

2

0

S

S

dg

wY

(34)

and so by calculating X we can obtain Y from the diagram and further derivate thevelocity as

16.0

L

V

V

Le

r0 dgY

S

Sw

(35)

In the above expressions, w0is obtained in m/s, g is the gravitational acceleration (g =9.81 m/s), deis the equivalent diameter of the tray hole covered by the valve (de= 2hS)and Q is introduced in m3/h.

We have obtained the following results which belong to the stable operating domainfor this type of trays.

Parameter Rectification StrippingX 0.525 0.8Y 1.1 0.4w0(m/s) 0.37 0.22


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26

Column Diameter

Since the vapour flow equals

0

2

iV w

4

DQ

(36)

we can calculate the column diameter as

0

Vi

w

Q4D

(37)

We obtain in the two cases the following values

rectifying section: Di= 0.83 m

stripping section: Di= 1.01 m

and so we take the same value for the same column (in order to ease the design),namely

column diameter: Di= 1.0 m

Tray design

Knowing the column diameter, we can easily obtain the total area

S = 0.785 m2

Since the hole area was imposed to be 15% of the total area, this gives

Sr= 0.118 m2

Part of the total area is occupied by two down comers, each having an area of

Ssin1802

1Sd

(38)

Sd= 0.071 m2

and so the active area of the tray is

Sa= S 2 Sd= 0.643 m2

The number of valves on each tray results from the formula

4

d

Sz

2

0

r

(39)

which gives z = 35.5 valves/tray, which is approximated to

z = 35 valves/tray

The disposition of the valves on the tray is done according to a network of equilateraltriangles. We chose

the distance of the marginal valves to the column wall1

= 50 mm

the distance of the marginal valves to the weir plate 2 = 75 mm weir plate height hz= 40 mm


27/51


27

A scale graphical representation of the tray design is given in the Appendix.

Liquid Entrainment

The tray spacing is taken to be 0.3 m (considering the low velocities), which is the

minimum accepted. We must calculate the liquid entrainment at this specific length inorder to assure it doesnt exceed a certain threshold.We use the relation

n

m

0

H

wKCe (40)

where C is a coefficient which for our case has the value C = 3.6 10-3and K, m and ndepend on the shape of the valve. For disc-like valves, the values are: K = 6.9, m = 2.7and n = 3. With this, the value of the entrainment becomes

e = 0.18 %

for the rectifying region and

e = 0.05 %

for the stripping region, which is very low and thus the tray spacing is safely chose.

Pressure Drop per Tray

For the vapour flow to pass through the valves and the liquid layer above, it isnecessary for the pressure p1underneath the tray to be greater than the pressure p2above the tray. The resistance that opposes to the vapour pass equals with

21 ppp (41)

and is composed from the resistance of the dried tray and the resistance of the liquidalone, thus

LU ppp (42)

The liquid component may be approximated with the relation

5.0hhgp dzLL (43)

where hzis the height of the weir, hdis the height of the liquid above the weir and

is the level difference. In its turn, hdmay be deduced from the semi-empirical relation3/2

d

L21d

l

QKKh

(44)

where QL is introduced in m3/h and Ki are two constants (depending on the

characteristics of the down comer). K2 = 2.84 for a plane down comer, while K1 isobtained as a graphical dependence between QL, ldand Di.

The dry component can be calculated from the formula

V

2

1U

2

w

p (45)

where w1is the local velocity of the vapours inside the valve,


28/51


28

r

V1

S3600

Qw

(46)

and is the hydraulic resistance coefficient of the valve, = 3.6.

After performing the computations, we get the following results.

Parameter Rectification StrippingK1 1.02 1.125K2 2.84 2.84ld(m) 0.707 0.707hd(mm) 16.2 50.4hz(mm) 40 40w1(m/s) 1.72 1.48

Lp ( N/m2) 288 462.5

Up ( N/m2) 75 60.5

p ( N/m2) 363 523

Later, these values are to be multiplied with the real number of trays in order toobtain the total pressure drop through the column.

We plan now to calculate the real number of trays based on the kinetics of the masstransfer. For this we need to calculate the diffusion coefficients, the partial masstransfer coefficients, the total mass transfer coefficient, and using this to draw thekinetic curve, (i.e. the real equilibrium curve) and to apply the McCabe-Thiele methodfor it, yielding the real number of trays.

Diffusion Coefficients

The values of the diffusion coefficients for the vapour and the liquid phase aredependent upon the properties of the diffusing compound and the diffusion medium.They can be calculated using several empirical relations.

For the vapour phase we have chosen the Maxwell relation modified by Galliland

2123/123/11

2/37

VM

1

M

1

VVp

T103.4D

(47)

where V are the molar volumes of the two components at their boiling points(computed as a sum of the molar volumes for each of the compounding elements), Mare the molar masses and the pressure p and the temperature T are introduced in atmand Kelvin, respectively.

For the liquid phase we have used the Wilke and Chang relationship

6.0

0

5.0

14

LV

TM104.7D

(48)


29/51


29

in which denotes a parameter for association (in our case, the mixture containingonly hydrocarbons, = 0), is the dynamic viscosity in P, M is the average molecularmass while V0is a parameter related to the solvent (in our case V0= 22.8).

We have obtained the following results.

Parameter Rectification StrippingM1(kg/kmole) 30 30M2(kg/kmole) 44 44M (kg/kmole) 35.7 35.7V1(cm

3/mole) 51.8 51.8V2(cm

3/mole) 74.0 74.0P (atm) 7.89 7.89

L (P) 1.10810-3 1.12810-3

T (K) 248.95 274.35DV(m

2/s) 8.0710-7 9.3310-7

DL(m

2

/s) 1.5210

-8

1.7810

-8

Partial Mass Transfer Coefficients

The equation for the mass transfer inside a phase is

ckN (49)

where N is the transported mass flow (kmole/h), k is the partial mass transfercoefficient (kmole/m2h c) and c is the concentration difference. Thus, thedimensions of the coefficient k are dependent upon the dimensions in which the

concentration is expressed. For c = y (molar vapour fractions), k =m/h.

The coefficients k are computed from criteria relations. For the vapour filmcoefficients, we have used the equation

5.05.04 ScRe108Sh (50)

where Sh is the Sherwood criterion, Re is the Reynolds criterion, Sc is the Schmidtcriterion and is a specific criterion for the hydraulic resistance of the system.Replacing the respective criteria with their expressions, we get

l

h

Dlw108Sh z

V

4

(51)

in which denotes the fraction of holes in the suspension (we took = 0.5) and lrepresents the specific dimension which we considered l = 1.

Knowing the Sh criterion, it is easy to calculate the partial mass transfer coefficientssince

V

V

D

l'kSh

(52)

The value we obtained for k is however related to the hole area S r and it must beconnected to the total area, thus


30/51


30

S

S'k'kk rVVV (53)

Finally a conversion from the unit m/s to kmole/m2h y must be performed, and for this

T

15.273

41.22

1kk

Vy

(54)

The same equation (50) is used to calculate the partial coefficients for the liquid film,only that here

1

lwRe (55)

and kLis already related to the total tray area, so no (53) conversion should be used.The values obtained after performing the calculations are gathered in a table below.

Parameter Rectification Stripping

w (m/s) 0.37 0.22

V (kg/m3) 14.18 15.31

L (kg/m3) 504.6 510.3

V (P) 8.0510-5 8.3910-5

L (P) 1.10810-3 1.12810-3

ReV 130 350 80 272ReL 336 764 199 053ScV 7.035 5.875ScL 144.46 124.53DV(m

2/s) 8.0710-7 9.3310-7

DL(m

2

/s) 1.5210

-8

1.7810

-8

ShV 3.064106 2.198106

ShL 2.232107 1.593107

kV(m/h) 8.901103 7.381103

kV(m/h) 1335 1107kL(m/h) 1221.35 1018.07

ky(kmole/m2h y) 65.3 49.1

kx(kmole/m2h x) 17 272.78 12 530.66

Kinetic Curve

We can now obtain, for several values of x between xD

and xW

, the values of the totalmass transfer coefficient as

xy k

m

k

1

1K

(56)

where m is the slope of the equilibrium line in each of these points, computed fromFenske equation like

2x11m

(57)

Now using the relation


31/51


31

V

a

n

*

n

1n

*

n

Q

SK

BC

ABln

yy

yylnN

(58)

we can obtain the coordinates of the points ynbelonging to the kinetic curve. The restof the y terms denote: yn* the equilibrium curve and yn+1 the operating lines. The

computations were done in the Excel sheet presented below.

The kinetic curve obtained is presented in Figure 10. Applying again the McCabe-Thiele method, this time between the operating lines and the newly drawn kineticcurve, we can obtain the real number of trays. We found

NRT = 10

feed tray 5

Thus the total pressure drop in the column is obtained as the sum between thepressure drop per tray and the real number of trays in the specified section of the

column (rectifying or stripping). We obtainp = 1815 Pa (for rectifying section)

p = 2615 Pa (for stripping section)

p = 4.43 kPa (for the total column)

This concludes our hydrodynamic calculus for the binary column. Further, we arerequired to calculate the height of the column and to draw a scale representationof it.

stripping feed rectifying

x n 0 0.05 0.1 0.15 0.2 0.25 0.3 0.35 0.4 0.45

ky (kmol/m2*h*Dy) 49.1 49.1 49.1 49.1 49.1 49.1 49.1 49.1 65.3 65.3

kx (kmol/m2*h*Dx) 12530.66 12530.66 12530.66 12530.66 12530.66 12530.66 12530.66 12530.66 17272.78 17272.78

m 4.84 3.406378 2.526813 1.948646 1.548392 1.2598917 1.045107 0.880907 0.658751 0.560624

Ky (kmol/m2*h*Dy) 48.18615 48.45327 48.61863 48.72794 48.8039 48.858797 48.89975 48.9311 65.13778 65.16189

V (kmol/h) 255.8797 255.8797 255.8797 255.8797 255.8797 255.87966 255.8797 255.8797 324.7585 324.7585

Nv 0.121087 0.121758 0.122174 0.122448 0.122639 0.1227773 0.12288 0.122959 0.128968 0.129016

BC/AB 0.885957 0.885362 0.884995 0.884752 0.884583 0.8844606 0.88437 0.8843 0.879002 0.87896

y*n 0 0.22555 0.380743 0.494058 0.580429 0.6484467 0.703398 0.748719 0.786737 0.819085

y n+1 -0.037873 0.068578 0.175028 0.281479 0.387929 0.4943798 0.60083 0.707281 0.729603 0.749922

AB 0.037873 0.156973 0.205715 0.212579 0.1925 0.1540668 0.102568 0.041438 0.057134 0.069163

BC 0.033554 0.138978 0.182057 0.18808 0.170282 0.1362661 0.090708 0.036644 0.050221 0.060791

AC 0.004319 0.017995 0.023658 0.024499 0.022218 0.0178008 0.01186 0.004794 0.006913 0.008371

y n -0.004319 0.207556 0.357085 0.469558 0.558212 0.6306459 0.691538 0.743925 0.779823 0.810713

x n 0.5 0.55 0.6 0.65 0.7 0.75 0.8 0.85 0.9 0.95 1

ky (kmol/m2*h*Dy) 65.3 65.3 65.3 65.3 65.3 65.3 65.3 65.3 65.3 65.3 65.3

kx (kmol/m2*h*Dx) 17272.78 17272.78 17272.78 17272.78 17272.78 17272.78 17272.78 17272.78 17272.78 17272.78 17272.78

m 0.482894 0.420276 0.369093 0.326721 0.291247 0.2612526 0.235664 0.213658 0.194597 0.177977 0.163399

Ky (kmol/m2*h*Dy) 65.18101 65.19641 65.20901 65.21944 65.22818 65.235569 65.24187 65.2473 65.252 65.25609 65.25969

V (kmol/h) 324.7585 324.7585 324.7585 324.7585 324.7585 324.75847 324.7585 324.7585 324.7585 324.7585 324.7585

Nv 0.129054 0.129085 0.129109 0.12913 0.129147 0.1291621 0.129175 0.129185 0.129195 0.129203 0.12921

BC/AB 0.878926 0.8789 0.878878 0.87886 0.878844 0.8788315 0.878821 0.878811 0.878803 0.878796 0.87879

y*n 0.846944 0.871188 0.892477 0.911321 0.928118 0.943184 0.956774 0.969095 0.980316 0.990578 1

y n+1 0.770242 0.790561 0.810881 0.831201 0.85152 0.87184 0.89216 0.912479 0.932799 0.953118 0.973438

AB 0.076702 0.080626 0.081596 0.08012 0.076597 0.071344 0.064614 0.056615 0.047517 0.03746 0.026562

BC 0.067416 0.070862 0.071713 0.070414 0.067317 0.0626994 0.056784 0.049754 0.041758 0.03292 0.023342

AC 0.009287 0.009764 0.009883 0.009706 0.00928 0.0086446 0.00783 0.006861 0.005759 0.00454 0.00322

y n 0.837657 0.861424 0.882594 0.901615 0.918837 0.9345393 0.948944 0.962233 0.974557 0.986038 0.99678


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32

The height of the column

The total height of the column is computed by the approximate formula

Ht Ha+ HV+ HB+ HS (59)

where HVis the distance from the first tray to the top of the column, HBis the distancefrom the last tray to the bottom of the column, HSis the distance from the bottom ofthe column to the ground, and Hais the height of the active section of the column

Ha= (NRT-1)H + NRT S (60)

in which H is the distance between the trays and S is the tray thickness. Note that

since the reboiler works at equilibrium, the trays to be actually considered in equation

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

x

y

Figure 10. The kinetic curve and the real number of trays


33/51


33

(60) are one less (the reboiler acts as a tray but is not computed in the total height ofthe column).

We have considered

H = 0.3 m

S = 5 mm

HV= 2H = 0.6 m

HB= 3H = 0.9 m

HS= 1.5 m

and thus the total height of the column reaches the value

Ht= 5.5 m

With this the binary column calculations are completed.

References

1. A. Bologa, T. Danciu Unit Operations Project Guide, Politehnica University ofBucharest, 2004.

2. A. Woinaroschy Unit Operations in Chemical Engineering, PolitehnicaUniversity Press, Bucharest, 1994.

3. O. Floarea, G. Jinescu, C. Balaban, P. Vasilescu, R. Dima Operatii si Utilaje inIndustria Chimica. Probleme pentru subingineri, Editura Didactica siPedagogica, Bucuresti, 1980.

4. * * * Ingineria Proceselor Fizice si Chimice. Ghid de proiect, PolitehnicaUniversity Press, Bucharest.


34/51

Appendixes


35/51


36/51

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Case Name: E:\MYX\SCHOOL\OU PROIECT\WORK\SIMULARE.HSC

Unit Set: SI

Date/Time: Thu Jun 03 19:29:47 2004

Pure Component: Ethane

Identification

Family / Class Chemical Formula ID Number Group Name CAS Number

Hydrocarbon C2H6 2 ---

UNIFAC Structure

(CH3)2

User ID Tags

Tag Number Tag Text

Critical/Base Properties

Base Properties Critical Properties

Molecular Weight

Normal Boiling Pt (C)

Std Liq Density (kg/m3)

30.07

-88.60

355.7

Temperature (C)

Pressure (kPa)

Volume (m3/kgmole)

Acentricity

32.28

4884

0.1480

9.860e-002

Temperature Dependent Properties

Vapour Enthalpy Vapour Pressure Gibbs Free Energy

Minimum Temperature (C) -270.0 -140.1 25.00

Maximum Temperature (C) 5000 32.25 426.9

Coefficient Name

a

b

cd

e

f

g

h

i

j

IdealH Coefficient

-1.768

1.143

-3.236e-0044.243e-006

-3.393e-009

8.821e-013

1.000

0.0000

0.0000

0.0000

Antoine Coefficient

44.01

-2569

0.0000-4.976

1.464e-005

2.000

0.0000

0.0000

0.0000

0.0000

Gibbs Free Coefficient

-8.579e+004

168.6

2.685e-0020.0000

0.0000

---

---

---

---

---

Additional Point Properties

Thermodynamic and Physical Properties Property Package Molecular Properties

Dipole Moment

Radius of Gyration

COSTALD (SRK) Acentricity

COSTALD Volume (m3/kgmole)

Viscosity Coefficient A

Viscosity Coefficient B

Cavett Heat of Vap Coeff A

Cavett Heat of Vap Coeff B

Heat of Formation (25C) (kJ/kgmole)

Heat of Combustion (25C) (kJ/kgmole)

Enthalpy Basis Offset (kJ/kgmole)

0.0000

1.826

9.830e-002

0.1458

7.231e-002

4.698e-002

0.2833

---

-8.474e+004

-1.429e+006

-9.670e+004

PRSV - Kappa

KD Group Parameter

ZJ EOS Parameter

GS/CS - Solubility Parameter

GS/CS - Molar Volume (m3/kgmole)

GS/CS - Acentricity

UNIQUAC - R

UNIQUAC - Q

Wilson Molar Volume (m3/kgmole)

CN Solubility

CN Molar Volume (m3/kgmole)

1.343e-002

2.198

0.0000

6.050

6.700e-002

0.1064

1.802

1.696

8.454e-002

4.287

9.798e-003

Hyprotech Ltd. HYSYS v3.0.1 (Build 4602) Page 1 of 1

Licensed to: TEAM LND


37/51

TEAM LND

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CANADA


Unit Set: SI


Pure Component: Propane

Identification



UNIFAC Structure

(CH3)2 CH2

User ID Tags

Tag Number Tag Text



Molecular Weight



44.10

-42.10

506.7

Temperature (C)

Pressure (kPa)

Volume (m3/kgmole)

Acentricity

96.75

4257

0.2000

0.1524





Coefficient Name

a

b

cd

e

f

g

h

i

j

IdealH Coefficient

39.49

0.3950

2.114e-0033.965e-007

-6.672e-010

1.679e-013

1.000

0.0000

0.0000

0.0000

Antoine Coefficient

52.38

-3491

0.0000-6.109

1.119e-005

2.000

0.0000

0.0000

0.0000

0.0000


-1.055e+005

264.8

3.250e-0020.0000

0.0000

---

---

---

---

---



Dipole Moment

Radius of Gyration










0.0000

2.431

0.1532

0.2001

7.112e-002

-6.538e-002

0.2783

---

-1.039e+005

-2.045e+006

-1.194e+005

PRSV - Kappa

KD Group Parameter

ZJ EOS Parameter



GS/CS - Acentricity

UNIQUAC - R

UNIQUAC - Q


CN Solubility


3.163e-002

3.007

0.0000

6.400

8.400e-002

0.1538

2.477

2.236

8.703e-002

5.999

1.072e-002




38/51

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Unit Set: SI


Pure Component: n-Butane

Identification



UNIFAC Structure

(CH3)2 (CH2)2

User ID Tags

Tag Number Tag Text



Molecular Weight



58.12

-0.5020

583.2

Temperature (C)

Pressure (kPa)

Volume (m3/kgmole)

Acentricity

152.0

3797

0.2550

0.2010





Coefficient Name

a

b

cd

e

f

g

h

i

j

IdealH Coefficient

67.72

8.541e-003

3.277e-003-1.110e-006

1.766e-010

-6.399e-015

1.000

0.0000

0.0000

0.0000

Antoine Coefficient

66.94

-4604

0.0000-8.255

1.157e-005

2.000

0.0000

0.0000

0.0000

0.0000


-1.284e+005

360.5

3.826e-0020.0000

0.0000

---

---

---

---

---



Dipole Moment

Radius of Gyration










0.0000

2.886

0.2008

0.2544

0.1001

-5.969e-002

0.2747

---

-1.262e+005

-2.660e+006

-1.456e+005

PRSV - Kappa

KD Group Parameter

ZJ EOS Parameter



GS/CS - Acentricity

UNIQUAC - R

UNIQUAC - Q


CN Solubility


3.951e-002

4.027

0.0000

6.730

0.1014

0.1953

3.151

2.776

9.966e-002

6.713

1.277e-002




39/51

TEAM LND

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Unit Set: SI

Date/Time: Wed Jun 02 20:26:56 2004

Distillation: First Column @Main

CONNECTIONS

Inlet Stream

STREAM NAME Sta e FROM UNIT OPERATION

Qreb1

F1

Reboiler

6__Main TS

Outlet Stream

STREAM NAME Sta e TO UNIT OPERATION

Qcond1

D1

W1

Condenser

Condenser

Reboiler

VLV-100Valve

MONITOR

Specifications Summary

Specified Value Current Value Wt. Error Wt. Tol. Abs. Tol. Active Estimate Used

CompRecov 2 Light 0.9900 0.9899 -3.301e-005 1.000e-002 1.000e-003 On On On

CompRecov 3 Heavy 0.9800 0.9799 -3.221e-005 1.000e-002 1.000e-003 On On On

Distillate Rate 479.8 kgmole/h 481.7 kgmole/h 3.901e-003 1.000e-002 1.000 kgmole/h Off On Off

Reflux Ratio 6.000 2.841 -0.5265 1.000e-002 1.000e-002 Off On Off

PROPERTIES

Properties : F1

Overall Vapour Phase Liquid Phase

Vapour/Phase Fraction

Temperature: (C)

Pressure: (kPa)

Molar Flow (kgmole/h)

Mass Flow (kg/h)

Std Ideal Liq Vol Flow (m3/h)

Molar Enthalpy (kJ/kgmole)

Mass Enthalpy (kJ/kg)

Molar Entropy (kJ/kgmole-C)

Mass Entropy (kJ/kg-C)

Heat Flow (kJ/h)

Molar Density (kgmole/m3)

Mass Density (kg/m3)

Std Ideal Liq Mass Density (kg/m3)

Liq Mass Density @Std Cond (kg/m3)

Molar Heat Capacity (kJ/kgmole-C)

Mass Heat Capacity (kJ/kg-C)

Thermal C onductivity (W/m-K)

Viscosity (cP)

Surface Tension (dyne/cm)

Molecular Weight

Z Factor

0.5500

59.67

1400

950.0

4.589e+004

88.21

-1.171e+005

-2423

123.3

2.553

-1.112e+008

1.093

52.81

520.3

537.4

122.4

2.534

---

---

---

48.31

---

0.5500

59.67

1400

522.5

2.359e+004

47.52

-1.053e+005

-2333

152.0

3.367

-5.503e+007

0.6341

28.62

496.4

517.9

95.08

2.106

2.302e-002

9.945e-003

---

45.14

0.7979

0.4500

59.67

1400

427.5

2.230e+004

40.69

-1.314e+005

-2519

88.33

1.693

-5.619e+007

9.500

495.7

548.2

557.9

155.8

2.986

7.792e-002

0.1012

6.136

52.17

5.326e-002

Properties : D1


Vapour/Phase Fraction 0.0000 0.0000 1.0000




40/51


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Unit Set: SI


Distillation: First Column @Main (continued)

COLUMN PROFILES

Reflux Ratio: 2.841 Reboil Ratio: 2.705 The Flows Option is Selected Flow Basis: Molar

Column Profiles Flows

Condenser

1__Main TS

2__Main TS

3__Main TS

4__Main TS

5__Main TS

6__Main TS7__Main TS

8__Main TS

9__Main TS

10__Main TS

11__Main TS

12__Main TS

13__Main TS

14__Main TS

15__Main TS

16__Main TS

Reboiler

Temperature (C)

9.259

26.84

37.52

45.61

53.29

60.15

65.3472.92

78.29

82.47

85.83

88.47

90.48

91.96

93.02

93.76

94.27

94.61

Pressure (kPa)

1400

1400

1400

1400

1400

1400

14001400

1400

1400

1400

1400

1400

1400

1400

1400

1400

1400

Net Liq (kgmole/h)

1368

1329

1289

1234

1188

1163

15991636

1659

1675

1690

1702

1713

1721

1727

1732

1735

---

Net Vap (kgmole/h)

---

1850

1811

1770

1716

1670

16451130

1168

1191

1207

1221

1234

1244

1253

1259

1263

1267

Net Feed (kgmole/h)

---

---

---

---

---

---

950.0---

---

---

---

---

---

---

---

---

---

---

Net Draws (kgmole/h)

481.7

---

---

---

---

---

------

---

---

---

---

---

---

---

---

---

468.3

Column Profiles Energy

Condenser

1__Main TS

2__Main TS

3__Main TS

4__Main TS

5__Main TS

6__Main TS

7__Main TS

8__Main TS

9__Main TS

10__Main TS

11__Main TS

12__Main TS

13__Main TS

14__Main TS

15__Main TS

16__Main TS

Reboiler

Temperature (C)

9.259

26.84

37.52

45.61

53.29

60.15

65.34

72.92

78.29

82.47

85.83

88.47

90.48

91.96

93.02

93.76

94.27

94.61

Liquid Enthalpy (kJ/kgmole)

-1.125e+005

-1.174e+005

-1.207e+005

-1.239e+005

-1.271e+005

-1.296e+005

-1.313e+005

-1.324e+005

-1.334e+005

-1.342e+005

-1.349e+005

-1.354e+005

-1.358e+005

-1.360e+005

-1.362e+005

-1.363e+005

-1.364e+005

-1.365e+005

Vapour Enthalpy (kJ/kgmole)

-9.317e+004

-9.813e+004

-1.014e+005

-1.034e+005

-1.052e+005

-1.069e+005

-1.084e+005

-1.117e+005

-1.139e+005

-1.156e+005

-1.170e+005

-1.181e+005

-1.189e+005

-1.196e+005

-1.201e+005

-1.204e+005

-1.206e+005

-1.208e+005

Heat Loss (kJ/h)

---

---

---

---

---

---

---

---

---

---

---

---

---

---

---

---

---

---




42/51

TEAM LND

Calgary, Alberta

CANADA


Unit Set: SI


0 2 4 6 8 10 12 14 16 180.000

10.0

20.0

30.0

40.0

50.0

60.0

70.0

80.0

90.0

100

Tempera

ture(C)

Temperature vs. Tray Number

HYSYS Column Profiles Specsheet:Column Temperature Profile

Column Temperature Profile

Column Stage

Condenser

1__Main TS

2__Main TS

3__Main TS

4__Main TS

5__Main TS

6__Main TS

7__Main TS

8__Main TS

9__Main TS

10__Main TS

11__Main TS

12__Main TS

13__Main TS

14__Main TS

15__Main TS

16__Main TS

Reboiler

Temperature

(C)

9.259

26.84

37.52

45.61

53.29

60.15

65.34

72.92

78.29

82.47

85.83

88.47

90.48

91.96

93.02

93.76

94.27

94.61




43/51

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Unit Set: SI


HYSYS Column Profiles Specsheet:

Column Properties Profile

Options Selected

Mass basis is selected

Stage

Condenser

1__Main TS

2__Main TS

3__Main TS

4__Main TS

5__Main TS

6__Main TS

7__Main TS

8__Main TS

9__Main TS

10__Main TS

11__Main TS

12__Main TS

13__Main TS

14__Main TS

15__Main TS

16__Main TS

Reboiler

Surface Tension

6.250 dyne/cm

5.921 dyne/cm

5.777 dyne/cm

5.760 dyne/cm

5.757 dyne/cm

5.713 dyne/cm

5.646 dyne/cm

5.271 dyne/cm

5.054 dyne/cm

4.910 dyne/cm

4.801 dyne/cm

4.715 dyne/cm

4.649 dyne/cm

4.598 dyne/cm

4.561 dyne/cm

4.534 dyne/cm

4.515 dyne/cm

4.503 dyne/cm

Mole Weight (Vap)

34.60

38.84

41.60

43.38

44.95

46.44

47.69

50.39

52.24

53.65

54.79

55.71

56.43

56.96

57.35

57.63

57.82

57.95

Density (Vap)

25.75 kg/m3

27.53 kg/m3

28.72 kg/m3

29.22 kg/m3

29.53 kg/m3

29.90 kg/m3

30.28 kg/m3

31.76 kg/m3

32.72 kg/m3

33.44 kg/m3

34.03 kg/m3

34.50 kg/m3

34.88 kg/m3

35.17 kg/m3

35.39 kg/m3

35.54 kg/m3

35.65 kg/m3

35.72 kg/m3

Viscosity (Vap)

9.188e-003 cP

9.407e-003 cP

9.514e-003 cP

9.622e-003 cP

9.737e-003 cP

9.836e-003 cP

9.902e-003 cP

9.915e-003 cP

9.933e-003 cP

9.951e-003 cP

9.966e-003 cP

9.977e-003 cP

9.985e-003 cP

9.989e-003 cP

9.992e-003 cP

9.994e-003 cP

9.996e-003 cP

9.996e-003 cP

Therm Cond (Vap)

2.020e-002 W/m-K

2.094e-002 W/m-K

2.140e-002 W/m-K

2.185e-002 W/m-K

2.234e-002 W/m-K

2.277e-002 W/m-K

2.308e-002 W/m-K

2.335e-002 W/m-K

2.358e-002 W/m-K

2.377e-002 W/m-K

2.392e-002 W/m-K

2.405e-002 W/m-K

2.414e-002 W/m-K

2.421e-002 W/m-K

2.425e-002 W/m-K

2.429e-002 W/m-K

2.431e-002 W/m-K

2.432e-002 W/m-K

Heat Cap (Vap)

1.969 kJ/kg-C

2.016 kJ/kg-C

2.051 kJ/kg-C

2.075 kJ/kg-C

2.099 kJ/kg-C

2.122 kJ/kg-C

2.142 kJ/kg-C

2.184 kJ/kg-C

2.213 kJ/kg-C

2.236 kJ/kg-C

2.255 kJ/kg-C

2.271 kJ/kg-C

2.283 kJ/kg-C

2.292 kJ/kg-C

2.299 kJ/kg-C

2.303 kJ/kg-C

2.307 kJ/kg-C

2.309 kJ/kg-C

Mole Weight (Lt Liq)

38.84

42.60

45.08

47.33

49.52

51.35

52.63

53.90

54.89

55.70

56.36

56.87

57.26

57.54

57.74

57.88

57.97

58.04

Stage

Condenser

1__Main TS

2__Main TS

3__Main TS

4__Main TS

5__Main TS

6__Main TS

7__Main TS

8__Main TS

9__Main TS

10__Main TS

11__Main TS

12__Main TS

13__Main TS

14__Main TS

15__Main TS

16__Main TS

Reboiler

Surface Tension

6.250 dyne/cm

5.921 dyne/cm

5.777 dyne/cm

5.760 dyne/cm

5.757 dyne/cm

5.713 dyne/cm

5.646 dyne/cm

5.271 dyne/cm

5.054 dyne/cm

4.910 dyne/cm

4.801 dyne/cm

4.715 dyne/cm

4.649 dyne/cm

4.598 dyne/cm

4.561 dyne/cm

4.534 dyne/cm

4.515 dyne/cm

4.503 dyne/cm

Density (Lt Liq)

478.1 kg/m3

479.3 kg/m3

480.6 kg/m3

484.2 kg/m3

487.5 kg/m3

489.0 kg/m3

489.2 kg/m3

485.0 kg/m3

482.6 kg/m3

481.0 kg/m3

479.6 kg/m3

478.5 kg/m3

477.6 kg/m3

476.9 kg/m3

476.4 kg/m3

476.1 kg/m3

475.8 kg/m3

475.7 kg/m3

Viscosity (Lt Liq)

9.235e-002 cP

9.180e-002 cP

9.165e-002 cP

9.389e-002 cP

9.592e-002 cP

9.717e-002 cP

9.723e-002 cP

9.479e-002 cP

9.349e-002 cP

9.267e-002 cP

9.207e-002 cP

9.160e-002 cP

9.126e-002 cP

9.101e-002 cP

9.083e-002 cP

9.072e-002 cP

9.064e-002 cP

9.060e-002 cP

Therm Cond (Lt Liq)

9.416e-002 W/m-K

8.837e-002 W/m-K

8.523e-002 W/m-K

8.251e-002 W/m-K

7.982e-002 W/m-K

7.759e-002 W/m-K

7.602e-002 W/m-K

7.345e-002 W/m-K

7.168e-002 W/m-K

7.039e-002 W/m-K

6.939e-002 W/m-K

6.862e-002 W/m-K

6.805e-002 W/m-K

6.763e-002 W/m-K

6.732e-002 W/m-K

6.711e-002 W/m-K

6.697e-002 W/m-K

6.687e-002 W/m-K

Heat Cap (Lt Liq)

2.983 kJ/kg-C

3.020 kJ/kg-C

3.039 kJ/kg-C

3.033 kJ/kg-C

3.029 kJ/kg-C

3.039 kJ/kg-C

3.054 kJ/kg-C

3.114 kJ/kg-C

3.153 kJ/kg-C

3.183 kJ/kg-C

3.207 kJ/kg-C

3.227 kJ/kg-C

3.243 kJ/kg-C

3.255 kJ/kg-C

3.264 kJ/kg-C

3.271 kJ/kg-C

3.275 kJ/kg-C

3.278 kJ/kg-C




44/51

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Calgary, Alberta

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Unit Set: SI


0 2 4 6 8 10 12 14 16 180.000

1.00e-001

0.200

0.300

0.400

0.500

0.600

0.700

0.800

0.900

1.00

MoleFraction

Ethane (Vap)

Propane (Vap)

n-Butane (Vap)

Ethane (Light)

Propane (Light)

n-Butane (Light)

Composition vs. Tray NumberHYSYS Column Profiles Specsheet:

Column Composition Profile

Options Selected

Fraction is selected as the composition basis Net is selected as flow basis

Molar basis is selected

Stage

Condenser

1__Main TS

2__Main TS

3__Main TS

4__Main TS

5__Main TS

6__Main TS

7__Main TS

8__Main TS

9__Main TS

10__Main TS

11__Main TS

12__Main TS

13__Main TS

14__Main TS

15__Main TS

16__Main TS

Reboiler

Ethane (Vap)

0.6808

0.3945

0.2355

0.1747

0.1552

0.1494

0.1471

0.0589

0.0221

0.0079

0.0027

0.0009

0.0003

0.0001

0.0000

0.0000

0.0000

0.0000

Propane (Vap)

0.3158

0.5857

0.7068

0.7018

0.6291

0.5340

0.4499

0.4336

0.3756

0.3032

0.2320

0.1701

0.1203

0.0825

0.0549

0.0353

0.0218

0.0124

n-Butane (Vap)

0.0034

0.0198

0.0577

0.1236

0.2157

0.3166

0.4031

0.5075

0.6023

0.6889

0.7653

0.8290

0.8794

0.9174

0.9451

0.9646

0.9782

0.9876

Ethane (Light Liq)

0.3945

0.1779

0.0925

0.0619

0.0500

0.0446

0.0417

0.0158

0.0057

0.0020

0.0007

0.0002

0.0001

0.0000

0.0000

0.0000

0.0000

0.0000

Propane (Light Liq)

0.5857

0.7506

0.7451

0.6460

0.5131

0.3936

0.3084

0.2699

0.2193

0.1689

0.1246

0.0889

0.0616

0.0416

0.0274

0.0175

0.0107

0.0061

n-Butane (Light Liq)

0.0198

0.0715

0.1623

0.2922

0.4369

0.5618

0.6500

0.7144

0.7750

0.8292

0.8747

0.9109

0.9383

0.9584

0.9726

0.9825

0.9893

0.9939




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Unit Set: SI


Distillation: Second Column (binary) @Main

CONNECTIONS

Inlet Stream

STREAM NAME Sta e FROM UNIT OPERATION

Qcond22

F22

Reboiler

4__Main TS

Outlet Stream

STREAM NAME Sta e TO UNIT OPERATION

Qreb22

D22

W22

Condenser

Condenser

Reboiler

MONITOR

Specifications Summary

Specified Value Current Value Wt. Error Wt. Tol. Abs. Tol. Active Estimate Used

Reflux Ratio 0.6846 1.432 1.092 1.000e-002 1.000e-002 Off On Off

Distillate Rate --- 192.8 kgmole/h --- 1.000e-002 1.000 kgmole/h Off On Off

Reflux Rate --- 276.1 kgmole/h --- 1.000e-002 1.000 kgmole/h Off On Off

Btms Prod Rate --- 288.9 kgmole/h --- 1.000e-002 1.000 kgmole/h Off On Off

CompRecov 1 0.9500 0.9499 -5.749e-005 1.000e-002 1.000e-003 On On On

CompRecov 2 0.9700 0.9700 -4.125e-006 1.000e-002 1.000e-003 On On On

PROPERTIES

Properties : F22


Vapour/Phase Fraction

Temperature: (C)

Pressure: (kPa)

Molar Flow (kgmole/h)

Mass Flow (kg/h)

Std Ideal Liq Vol Flow (m3/h)

Molar Enthalpy (kJ/kgmole)

Mass Enthalpy (kJ/kg)

Molar Entropy (kJ/kgmole-C)

Mass Entropy (kJ/kg-C)

Heat Flow (kJ/h)

Molar Density (kgmole/m3)

Mass Density (kg/m3)

Std Ideal Liq Mass Density (kg/m3)

Liq Mass Density @Std Cond (kg/m3)

Molar Heat Capacity (kJ/kgmole-C)

Mass Heat Capacity (kJ/kg-C)

Thermal C onductivity (W/m-K)

Viscosity (cP)

Surface Tension (dyne/cm)

Molecular Weight

Z Factor

0.1433

-10.76

800.0

481.7

1.852e+004

41.44

-1.121e+005

-2915

106.4

2.767

-5.399e+007

2.469

94.93

447.0

464.7

98.59

2.564

---

---

---

38.45

---

0.1433

-10.76

800.0

69.02

2374

5.888

-9.348e+004

-2717

163.7

4.760

-6.451e+006

0.4221

14.52

403.3

423.2

60.26

1.752

1.741e-002

8.277e-003

---

34.40

0.8688

0.8567

-10.76

800.0

412.7

1.615e+004

35.55

-1.152e+005

-2944

96.82

2.474

-4.754e+007

13.07

511.3

454.2

470.6

105.0

2.683

0.1086

0.1146

8.901

39.13

2.806e-002




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TEAM LND

Calgary, Alberta

CANADA


Unit Set: SI


Distillation:

Andreiadis Design of a Rectification Unit

Documents

Transcript of Andreiadis Design of a Rectification Unit