Andreiadis Design of a Rectification Unit

Unit Operations Project

Politehnica University of Bucharest Faculty of Engineering in Foreign Languages Chemical Engineering Division

Design of a Rectification Unit

Eugen S. Andreiadis Group 1244 E Year 2003 - 2004

Unit Operations Project — Design of a rectification unit

Foreword

The objective of this project is the calculation of two rectification towers, one employing 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 problems which may arise in multi-component rectification. Following, the first column calculation is done, starting with mass balances and the determination of the temperature profile, and continuing with the minimum reflux ratio (by the Underwood method) and the total reflux ratio (employing the Fenske equation). An optimal reflux ratio and the corresponding theoretical number of trays are obtained by the Galliland method, after which all the previously computed values are verified against a Hysys simulation, thus finalizing the first column calculation. The second column is assumed to be fed with a binary mixture, simplifying in this way the computations. This more rigorous analysis also includes a hydrodynamic calculus and the determination of the mass transfer coefficients and the real number of trays. Knowing this specific details, it is possible to draw a scaled representation of the second column. The project ends with several Appendixes containing raw data extracted from Hysys concerning the first and also the second column simulations.

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Unit Operations Project — Design of a rectification unit

Project Outline

Project Brief 4 Signification of Symbols 5 1. Justification 6 2. 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 19 3.1. Theoretical Number of Trays 19 3.2. Hydrodynamic Calculus 22 References 33 Appendixes 34

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Unit Operations Project — Design of a rectification unit

Project Brief A ternary mixture containing the following components ƒcomponent 1: C2H6 (ethane), ƒcomponent 2: C3H8 (propane) and ƒcomponent 3: nC4H10 (n-butane) is to be separated by continuous rectification, given the conditions below: Parameter

Symbol

Value Column 1

Column 2

Feed flow Feed composition (mole fractions)

F xF,i

Thermal state of feed Degrees of recovery

q s

Pressure at the top of the columns Type of condenser Type of reboiler Type of column

p -

950 kmole/h xF,1 = 0.2 xF,2 = 0.3 0.45 sv = 0.99 sh = 0.98 14 bar total equilibrium plates

sv = 0.95 sh = 0.97 8 bar total equilibrium valves

Project content: 1. Justification of chosen methods and solutions. 2. Thermodynamic and physical properties. 3. First column calculation using short-cut methods: minimum and optimum reflux, number of theoretical plates. 4. Binary column calculation using rigorous methods: - material balance; - thermal state of feed; - optimum reflux and number of theoretical plates; - column diameter and hydraulic calculations; - mass transfer coefficients and real number of plates. 5. Drawings 6. References Separation scheme imposed:

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Unit Operations Project — Design of a rectification unit

Signification of Symbols

F — feed stream or feed flow D — distillate stream or distillate flow W — waste stream or waste flow yi — vapour composition of component i xi — liquid composition of component i Įij — relative volatility of i to j si — degrees of recovery LKC — light key component HKC — heavy key component KD,i — distribution coefficient of component i Pr — reduced pressure Pc — critical pressure Pv — vapour pressure Tr — reduced temperature q — thermal state of feed Lmin — 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 U — density K — dynamic viscosity V — surface tension w — vapour or liquid velocity Di — diameter of the column Sr — free area of the tray (hole area) S — transversal section area Sd — downcomer area Sa — active area z — number of valves hz — weir height H — tray spacing t — valve spacing ld — weir length de — equivalent diameter dS — valve tray diameter d0 — valve hole diameter hS — maximum valve height e — liquid entrainment G1 — valve mass G S — tray thickness DV, DL — diffusion coefficients kx, ky — partial mass transfer coefficients K — total mass transfer coefficient Re — Reynolds criterion Sc — Schmidt criterion Sh — Sherwood criterion

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Unit Operations Project — Design of a rectification unit

1. Justification Separation of individual substances in a homogeneous liquid mixture or complete fractionation of such mixtures into their components is an important step in many production 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 of intimate contact between the homogeneous mixture and a second phase, which thereby allows mass transfer to occur between phases. The thermodynamic conditions are chosen so that only the component to be separated enters the second phase. The phases 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: ƒCreation of a two-phase system ƒMass transfer between phases ƒSeparation of the phases A large number of separation techniques utilize this very effective principle or modifications thereof. Absorption, desorption, evaporation, condensation, and distillation involve a gaseous and a liquid phase; solvent extraction uses two liquid phases. Separating techniques that utilize a fluid phase and a solid phase are adsorption, crystallization, drying, and leaching. In most of these separations, the necessary 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, in distillation, 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 pure substances. Indeed, distillation requires energy only in the form of heat, which can subsequently be removed from the system. The vapour and liquid are brought into intimate contact by countercurrent flow and mass exchange occurs because the two phases are not in thermodynamic equilibrium. The phases produced during rectification are formed by evaporation and condensation of the initial mixture. The separation of a liquid mixture into its pure components can be controlled solely via the heat supply. The basis of this unit operation is the volatility difference between the properties of the liquid and the vapour phase, respectively, on the vapour-liquid equilibrium of the system. Among the main methods used in distillation practice (differential, flash, batch, azeotropic, extractive distillation, a.s.o.), we are interested here solely in rectification. 6

Unit Operations Project — Design of a rectification unit

In differential or flash distillation, the vapour leaving the still at any time is in equilibrium with the remaining liquid, and there will normally be only a small increase in the concentration of the more volatile component. The essential merit of rectification 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 “fractionating column” which enables successive vaporisation and condensation to be accomplished in one unit.

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

Figure 1 and 2 illustrate the differences between continuous distillation and multiple distillation, i.e. rectification. The fractionating column consists of a cylindrical structure divided into sections by a series of perforated trays which permit the upward flow of vapour. The liquid reflux flows across each Figure 1. Continuous distillation of trays over a weir and downcomer to the tray a binary mixture a-b below. The vapour rising from the top tray A) Flow diagram showing symbols for total molar streams or flow rates (F, L, etc.) and passes to a condenser and then to some form mole fractions of the more volatile component of reflux divider where part is withdraw as a a (x , x , etc.); B) A y-x diagram showing F D equilibrium and operating lines product D and the remainder is returned to the top tray as reflux. The reflux stream is frequently passed from the condenser through a reflux drum and then pumped to the column at a rate determined by a suitable control device. The liquid in the base of the column is heated, either by condensing steam or by a hot oil steam, and the vapour rises through the perforations to the bottom tray. A more commonly used arrangement consists of an external reboiler. Here the liquid 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. 7

Unit Operations Project — Design of a rectification unit

Vapour of composition yw enters the bottom tray (say n) where it is partially condensed and then revaporised to give vapour of composition yn. This operation of partial condensation and partial vaporisation of the reflux liquid id repeated on each tray. The feed stream is introduced on some intermediate tray where the liquid has approximately the same composition as the feed. The part of the column above the feed point is known as the rectifying section, while the lower part is known as the stripping section. Figure 3. Flow of vapour and liquid through a rectification column

Multi-component mixture rectification is a more frequently operation in industrial chemistry 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 an approach exhibits is the optimal selection and sub-sequencing of separation operations and adequate equipment in order to meet the specified process requirements. For example, let us consider such a mixture of 3 components A, B and C (given in order of decreasing volatility) which should be separated until a given purity by continuous rectification. Three schemes can be imagined, as follows from Figure 4. A(+B)

A+B+C

A(+B)

1

A+B(+C)

A(+B)

2

2

(A+)B A+B+C

2

A+B+C

1

(A+)B(+C)

1 B(+C)

(A+)B+C

(B+)C

(B+)C

(B+)C 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 contains some 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. 8

Unit Operations Project — Design of a rectification unit

Scheme 2 uses the first column for obtaining C like waste (in the bottom). The distillate, also containing a small quantity of C, is separated further in the second column in order to obtain two fractions corresponding of A and B. According to the relative volatility, the component C cannot be found in the distillate of the second column. The third scheme is much more complicated, because of the side streams in both columns, but is also more economic. In practice, choosing one of the schemes is a matter of costs. The first scheme is more economic 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 to eliminate it first. Anyway, one column can reasonably isolate only one component, so for n components we have n-1 columns. The sequencing scheme imposed for our project is the second, so we are to separate component 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 important approximation that has to be made concerns the so called key components. The volatile components are called light and the less volatile ones are called heavy. The light 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 be neglected). If all light components are significant fractions of the waste, the lightest one will be called the light key component. The heavy key component (HKC) is that heavy component which could be found as an important amount in the distillate (or the heaviest one, if all key components are present in the distillate). LKC: HKC: external:

component 2 component 3 component 1

(propane) (butane) (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 (which will be verified later using the Shiras, Hanson and Gibson equation), the calculus evolves, with some modifications, like for a binary mixture composed of these key components.

2.1. Mass Balances A simple representation of the flows and the composition of the flows in the case of the first column is given below in Figure 5. The separation degrees are as follows 9

Unit Operations Project — Design of a rectification unit

sv

D ˜ x D2 F ˜ xF2

(1)

sh

W ˜ xW3 F ˜ x F3

(2)

V

L

F

D

and the key component approximation actually implies x W1

(3)

0

The mass balances are

F

V

(4)

DW

F ˜ x F1

D ˜ x D1  W ˜ x W 1

(5)

F ˜ x F2

D ˜ x D2  W ˜ x W 2

(6)

W L’ Figure 5. First Column

The system above contains only six equations but should calculate eight unknowns. We also have to add the stoichiometric restrictions

¦x ¦x

Dj

Wj

1

(7)

1

(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 Data for the mass balance: F, x Fj, s v, s h

s v (%) =

99 s h (%) =

98

Calculus hypothesis: x W1 = 0 Flows (kmol/h): F=

950

D= 481.65 W= 468.35 Total balance: 0

Concentrations x F=

1 0.2

2 0.3

x D= 0.3945 0.0000 x W= Partial balances 0

0.5858 0.0061 2E-14

3 Verification: 0.5 1 0.0197 0.9939 0

1 1

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

sv

D ˜ x D1 F ˜ x F1

(1’)

sh

W ˜ xW2 F ˜ x F2

(2’)

x D3

0

(3’)

F

DW

(4 )

F ˜ x F2

D ˜ x D2  W ˜ x W 2

(5’)

F ˜ x F3

D ˜ x D3  W ˜ x W 3

(6’) 10

Unit Operations Project — Design of a rectification unit

¦x ¦x

Dj

Wj

1

(7 )

1

(8 )

We get the values Data for the mass balance: F, xFj , s v , s h

s v (%) =

95 s h (%) =

97

Calculus hypothesis: xD3 = 0 Flows (kmol/h): F=

481.65

Concentrations: xF=

D= W= Total balance:

188.975 292.675 0

xD = xW = Partial balances:

1 0.3945

2 0.5858

0.95521 0.044792 0.03246 0.935119 0 0

3 Verification: 0.0197 1 0 0.03242 0

1 1

We see that in both cases the mass balances close, although for some columns there may 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 components’ distribution. 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 condenser employed. Thus, for the case of a total condenser, it means that the vapour fraction on the first theoretical stage has a known composition yDi = xDi. The pressure being given, 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 temperature is between the boiling point of the liquid fraction W and the temperature of the vapour which enters the N+1 stage. In order to simplify the calculus, the dew point of W is considered to be the searched temperature. The calculations are done employing a Pascal program which tries several temperatures until the needed conditions are fulfilled (verification is done using the Riedel — Plank — Miller relationship, see below). We found the values TD = 29.21 ˚C TW = 95.85 ˚C

(for the top of the column) (for the bottom of the column)

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Unit Operations Project — Design of a rectification unit

Verification of Computed Values The two values obtained have been verified using the De Priester diagram and also the 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 estimated values for Kdi verify the relation

yi

¦x ¦K

(9)

1

i

Di

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

¦ y ¦ x ˜K i

i

Di

(10)

1

The Riedel-Plank-Miller relationship states that 

log Pr,vi

>

Gi 3 1  Tr2,i  gi 1  Tr ,i Tr ,i

@

(11)

where Pr,vi is the reduced vapour pressure of component I, Tr,I is the reduced temperature and Gi and gi are two constants specific for each component. The relation allows the estimation of Pv similarly to the Antoine equation, only that it offers much 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 in the case of n-Butane being presented below

G  1.50014 g  1.81934 Tcr  425.12 lg  §¨

G·

© Tr ¹

lg

Pcr  37.34

˜ ª¬ 1  Tr  g ˜ ( 1  Tr)

1.135 lg

Pvap  10 ˜ Pcr Pvap Kd 

Kd

2.739 Pvap

Tr 

2

º¼

3

302.36 Tcr

temp redusa = temp / temp critica pres redusa = pres / pres critica 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 Parametrii Miller (G si g) C2H6 1.38881 1.55945 etan 1.44933 1.69639 propan C3H8 nC4H10 1.50014 1.81934 n-butan

13.82 0.198

from Lange's Handbook of Chemistry, 15th Ed, John A. Dean

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 the basis of the same equation. 12

Unit Operations Project — Design of a rectification unit Table 1. Riedel — Plank — Miller validation T (˚C) 29.21

95.85

Comp 1 2 3 1 2 3

KDi 3.268 0.751 0.198 12.213 2.988 0.996

xDi 0.1207 0.7800 0.0995 0.0000 0.0182 0.9899

¦x

Di

1.0023

1.0081

The results obtained using the De Priester diagram are also given in Table 2. The values obtained are also in good agreement with the expected ones. Table 2. De Priester validation T (˚C) 29.21

95.85

Comp 1 2 3 1 2 3

KDi 2.55 0.82 0.26 5.15 2.4 1.0

xDi 0.1547 0.7144 0.0758 0.0000 0.0146 0.9939

¦x

Di

0.9449

1.0085

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

q

FL F

(12)

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

¦x

* F,i

x F,i D,i (1  q)

¦ qK

1

(13)

where x F* ,i denotes the composition of the liquid part of the feed. We have obtained TF = 60.21 ˚C and the following phase composition

x F* ,1 = 0.0523

y F* ,1 = 0.3208

x F* ,2 = 0.2347

y F* ,2 = 0.3534

x F* ,3 = 0.7129

y F* ,3 = 0.3251

The average temperature in the column is computed as an arithmetic mean between the top and the bottom temperature, giving Tm = 62.53 ˚C 13

Unit Operations Project — Design of a rectification unit

2.3. Minimum Reflux Ratio The minimum reflux ratio requires an infinite number of theoretical plates in order to separate the key components as required, so it is a value that cannot be practically achieved, 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 temperature in the column, we can obtain the distribution coefficients KD,i and from here the relative volatilities (given in respect to the HKC) according to

Di

K D,i K D,HKC

(14)

The results are presented in Table 3. Table 3. Relative volatilities T (˚C) 62.53

Comp 1 2 3

KDi 1.579 6.412 0.483

Įi 13.275 3.269 1

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

x D,i ˜ D x F,i ˜ F

D i  1 x D,LKC ˜ D D LKC  D i x D,HKC ˜ D  D LKC  1 x F,LKC ˜ F D LKC  1 x F,HKC ˜ F

(16)

where the relative volatilities D i are calculated as above. If

x D,i ˜ D x ˜D < -0.01 or D,i > 1.01, x F,i ˜ F x F,i ˜ F then the component i is probably not distributed, but if

x D,i ˜ D  >0.01, 0.99@, x F,i ˜ F then i is distributed for sure (is not external). In our case the above formula is applied only once in order to verify if the first component is really external. We obtain

x D,i ˜ D = 5.268 > 1.01 x F,i ˜ F and so indeed component 1 is not distributed. 14

Unit Operations Project — Design of a rectification unit

Underwood Method The value of the minimum reflux ratio Lmin is only needed to verify the real reflux ratio and thus a very accurate method is not required. We chose the algorithm proposed by Underwood, which considers constant the relative volatilities D i and the ratio L/V and assures a reasonable value after a moderate effort. The following system has to be solved. ­ D i ˜ F ˜ x F,i °¦ D  ) ° i ® D ˜ i D ˜ x D,i ° ¦ °¯ Di  )

F1  q

(17-18) DL min  1

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 D LKC ) gives a variant of 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 ) was located 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 being returned in the column. Since the products’ distribution changes with the reflux ratio, the calculus at total reflux could give precious information regarding the final compositions. In order to find the minimum number of trays corresponding to the infinite reflux, we use an extended form of Fenske equation, written for key components,

Nmin  1

ª D ˜ x D,LKC W ˜ x W ,HKC º 1 ˜ log« ˜ » log DLKC ¬ D ˜ x D,HKC W ˜ x W ,LKC ¼

(19)

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

TD TW D LKC ˜ D LKC

(20)

TD TW = 3.793, D LKC = 3.000 We find, employing the same Riedel-Plank-Miller relation, D LKC and thus DLKC = 3.373, following that

Nmin = 6 Equation (19) can also be used to calculate the distribution of other components that the key ones in the case of total reflux. 15

Unit Operations Project — Design of a rectification unit

D ˜ x D,i W ˜ x W ,i

DLKC

Nmin 1

˜

D ˜ x D,HKC W ˜ x W ,HKC

(21)

In this case, only the external 1 component’s distribution could be calculated. This will give 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 value obtained 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 still possible to obtain non-zero values for some mass balance verifications. Data for the mass balance: F, x Fj, s v, s h

s v (%) =

Calculus hypothesis: D*x D1=alfa*W*x W1

a1 =

alfa =

99 s h (%) =

98

3.373 Nmin =

6

101.37

Flows (kmol/h): F=

950

D= 479.79 W= 470.21 Total balance: 0

Concentrations x F=

1 0.2

2 0.3

3 Verification: 0.5 1

x D= 0.3921 0.58806 0.0198 0.0039 0.00606 0.98999 x W= Partial balances -7E-15 2E-14 0

1 1

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 relative volatilities. The following correlation has been made,

Y

0.545827  0.591422 ˜ X  0.002743 ˜ X 1

NT  Nmin L  L min and X . NT  2 L 1 For several values of the reflux ratio L (L > Lmin), X is calculated with the above formula and Y from equation (22). From this, a graphical dependence of the type NT = f(L) can be obtained. The optimal reflux ratio is considered to be that value where the curve tends to remain constant (obtained by drawing two tangents to the curve and locating their interception point). The computations were done using an Excel sheet and the results obtained are plotted in Figure 6. A part of the data stream employed is given in the sheet aside.

(22)

where Y

L 1.02 1.50 2.00 2.50 3.00 3.50 4.00 4.50 5.00 5.50 6.00

X 0.009307 0.199520 0.332933 0.428229 0.499700 0.555289 0.599760 0.636145 0.666467 0.692123 0.714114

Y 0.835049 0.441574 0.357162 0.298969 0.255783 0.222357 0.195689 0.173908 0.155780 0.140453 0.127325

N 46.50 12.33 10.44 9.41 8.75 8.29 7.95 7.68 7.48 7.31 7.17

It follows that Lopt = 3 and NTT = 9.

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Unit Operations Project — Design of a rectification unit

Figure 6. The Gilliland method for obtaining the optimum reflux ratio

9.80

Nt

8.80

7.80

6.80

5.80 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

2.6 Hysys Simulation Following the determination of the optimum reflux ratio, we chose to verify all the above calculations using a recent version (3.0.1) of the powerful software package Hysys, the most famous process simulator. The chosen method was the so-called “short-cut distillation,” providing initial estimates for a much rigorous later calculus. This method is based on the Fenske-Underwood equations, and it gives Nmin and Lmin together with the number of ideal trays, the optimal feed location and other various properties for the components in the mixture. Procedure The procedure goes as follows: firstly, from the list of available components are chosen the ones to be separated. For each a set of physical and thermodynamical properties can be consulted. For reference, the properties found for our components are 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-Robinson equation of state). 17

Unit Operations Project — Design of a rectification unit

Now the access to the simulation manager is granted. This visual environment allows us to select and connect the equipments that are going to be used. Thus, we have introduced the two columns and joined them through a valve (since there exists a pressure difference between the two). We were required to name each flow rate and to 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 have introduced the previously determined temperatures at the two ends, but we have chosen not to do so and let the simulator compute them. The reflux ratio L as well as the liquid rate can also be specified in advance. Instead, we have introduced the imposed recovery rate sv and sh. More information are needed to fully characterise the feed stream F, for example the composition of the stream xFi, the pressure, the thermal state q and of course the molar flow rate. At this moment the simulator can calculate all the other characteristics of the streams and the column that we are interested in. Data inputs and outputs F = 950 kmole/h xF1 = 0.2 xF2 = 0.3 P = 1400 kPa q = 0.45 (q’ = 0.55) Having introduced the above inputs, we were able to try next several values for the number of plates until the value for the reflux ratio approached the Lopt computed with the Galliland method. However, for the NTT = 9 determined at paragraph 2.5., the column didn’t converge. The convergence was assured only for values of NTT greater than 11. Moreover, the Lopt = 3 was reached only at much higher number of trays, as can be seen from the Table 4. Note that since the reboiler works at equilibrium, it is considered a theoretical tray also. The location of the feed point was chosen as to correspond to the same composition of the vapour-liquid system inside the column as in the feed stream. For this, the compositions of the liquid key components on each tray were extracted and then divided. Knowing the thermal state of the feed q

x LF,i x F,i

(23)

Table 4. First column NT Nfeed L 11 5 21.4 12 5 8.08 13 6 5.43 14 6 4.18 15 6 3.52 16 6 3.12 17 6 2.84 18 7 2.62 19 7 2.46 20 7 2.34

the compositions of the liquid key components in the feed can be obtained and their ratio calculated. This ratio is compared then with the ratio on each tray and the closest 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 in this conditions are presented in the Appendix. 18

Unit Operations Project — Design of a rectification unit

The same was applied for the second column, as we tried to find the ideal number of trays and the location of the feed point that would minimize the reflux ratio, in the same time keeping a smooth temperature profile in the column. The optimum results obtained 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 1 2

NT 17 9

Nfeed 6 4

L 2.84 1.43

3. Second Column Calculation If the first column was fed with a ternary mixture, this one is considered to be a binary column, in order to simplify the calculations. For this, the component 3 was distributed so that the compositions of the first two would total 1. Also, a new column was simulated in Hysys according to the conditions above. Employing the simulator, various physical and thermo-dynamical properties for the selected components (1 and 2) 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 several fields needed to be introduced in order to obtain the mass balance, Lmin, Nmin and NTT, as detailed in the case of the first column. The results are presented below. Components

Molar Masses

C2H6 etan C3H8 propan Technological Imputs F= 481.67 kmol / h xF = 0.402441 kmol 1/ kmol am xD = 0.955253 kmol 1/ kmol am xW = 0.033545 kmol 1/ kmol am p= 800 kPa q= 0.857

30 kg / kmol 44 kg / kmol

Antoine Constants A B 44.0103 -2568.82 52.3785 -3490.55

C 0 0

D -4.97635 -6.10875

Mass balance F= D= W=

481.67 192.7791 288.8909 0

E 1.46E-05 1.12E-05

F 2 2

Boiling Points x1 x2 0.402441 0.597559 0.955253 0.0447466 0.033545 0.9664548 -4.09E-14 0

1 1 1

-39.23 18.1

After obtaining the mass balance, the liquid-vapour diagram is to be plotted. For several 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 accurate predictions at higher pressures) was employed to give the vapour pressures for the two components. (Note: for the calculation of the ternary column, we have used instead the Riedel-Plank-Miller relation.) The equation is as follows 19

Unit Operations Project — Design of a rectification unit

ln PV

a

b  d ˜ ln T  e ˜ T f Tc

(24)

where the pressure is obtained in kPa, the temperature is introduced in K and the parameters a to f have to be obtained from tables (or the simulator). Knowing the vapour pressures, it is easy to calculate the corresponding molar fractions x and y with the formulas

x1

p  PV,2 PV,1  PV,2

(25)

y1

x1 ˜

P1 p

(26)

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

D1,2

PV,1 PV,2

(27)

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

Liquid-Vapour Equilibrium n 0 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20

T -39.23 -36.3635 -33.497 -30.6305 -27.764 -24.8975 -22.031 -19.1645 -16.298 -13.4315 -10.565 -7.6985 -4.832 -1.9655 0.901 3.7675 6.634 9.5005 12.367 15.2335 18.1

pV1 799.6177 876.7525 959.2057 1047.196 1140.946 1240.683 1346.638 1459.046 1578.149 1704.192 1837.427 1978.114 2126.515 2282.905 2447.562 2620.775 2802.841 2994.066 3194.766 3405.269 3625.912

pV2 115.0705 129.9149 146.1995 164.0144 183.4517 204.6055 227.5715 252.4472 279.3319 308.3261 339.5322 373.0538 408.9961 447.4658 488.5708 532.4204 579.1256 628.7986 681.5532 737.5046 796.7699

x 1.000558 0.89723 0.804177 0.720107 0.643918 0.574662 0.511523 0.453798 0.400879 0.352236 0.30741 0.266 0.227656 0.192071 0.158974 0.128129 0.099327 0.072381 0.04713 0.023426 0.001142

y 1.00008 0.983311 0.964213 0.942617 0.918345 0.891217 0.861046 0.827641 0.790808 0.750347 0.706054 0.657723 0.605143 0.548099 0.486374 0.419748 0.347996 0.270893 0.18821 0.099715 0.005175

alfa12 6.948936 6.748666 6.560936 6.384781 6.219327 6.063783 5.917429 5.779609 5.649728 5.527239 5.411645 5.302489 5.199353 5.101854 5.009637 4.92238 4.839781 4.761566 4.687479 4.617285 4.550764 5.533556

1 0.9 0.8 0.7 0.6 0.5 0.4 0.3 0.2 0.1 0 0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

Next the minimum reflux ratio is computed as

L min

x D  y F* y F*  x F

(28)

where yF* represents the equilibrium concentration corresponding to xF. The optimal reflux is computed with the Woinaroschy equation 20

Unit Operations Project — Design of a rectification unit

L opt

1  0.192 ˜

L min

Nmin L min

(29)

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

y  xD

L x  x D L 1

(30)

and for the stripping line

y  xW

L F x  x W L 1

(31)

while the Fenske equation is

y D  D  1 y

x

(32)

1

Minimum reflux ratio yf* = Lmin =

0.9

0.788436 kmol 1 / kmol am 0.432174

0.8 0.7

Minimum number of trays 0.6

n 0 1 2 3 4

x 0.955253 0.794151 0.41079 0.111895 0.022262

Nmin =

y 0.955253 0.794151 0.41079 0.111895 0.022262

4

0.5 0.4 0.3 0.2 0.1 0

Optimum reflux Lopt =

0.684615

QL = QV = QL ' = QV ' = xq =

131.9794 324.7585 544.7706 255.8797 0.351162

kmol / h kmol / h kmol / h kmol / h kmol 1 / kmol am

Theoretical number of trays

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

1 0.9 0.8 0.7

n cond 1 2 3 alim

4 5 6 7 8 9

x 0.955253 0.794151 0.593316 0.432244 0.342821 0.288771 0.19771 0.100884 0.03739 0.007808

y 0.955253 0.955253 0.889783 0.808165 0.742706 0.691997 0.576924 0.383053 0.176911 0.04173

0.6 0.5 0.4 0.3 0.2 0.1 0

NTT =

9

21

Unit Operations Project — Design of a rectification unit

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

Surface Tension Mole Weight (Vap) 8.480 dyne/cm 30.19 8.769 dyne/cm 30.7 9.001 dyne/cm 31.88 8.978 dyne/cm 33.47 8.874 dyne/cm 34.68 8.753 dyne/cm 35.78 8.556 dyne/cm 37.41 8.310 dyne/cm 39.36 8.075 dyne/cm 41.22 7.895 dyne/cm 42.66 Surface Tension Mole Weight (Lt Liq) 8.480 dyne/cm 30.7 8.769 dyne/cm 32.78 9.001 dyne/cm 35.78 8.978 dyne/cm 38.15 8.874 dyne/cm 39.38 8.753 dyne/cm 40.27 8.556 dyne/cm 41.32 8.310 dyne/cm 42.32 8.075 dyne/cm 43.1 7.895 dyne/cm 43.63

Density (Vap) 14.38 kg/m3 14.24 kg/m3 14.18 kg/m3 14.35 kg/m3 14.58 kg/m3 14.82 kg/m3 15.24 kg/m3 15.79 kg/m3 16.35 kg/m3 16.81 kg/m3 Density (Lt Liq) 480.6 kg/m3 492.6 kg/m3 504.9 kg/m3 510.3 kg/m3 511.4 kg/m3 511.4 kg/m3 510.5 kg/m3 508.7 kg/m3 506.6 kg/m3 504.9 kg/m3

Viscosity (Vap) 7.732e-003 cP 7.855e-003 cP 8.052e-003 cP 8.214e-003 cP 8.293e-003 cP 8.342e-003 cP 8.387e-003 cP 8.408e-003 cP 8.405e-003 cP 8.390e-003 cP Viscosity (Lt Liq) 9.783e-002 cP 0.1044 cP 0.1110 cP 0.1143 cP 0.1146 cP 0.1142 cP 0.1130 cP 0.1113 cP 0.1095 cP 0.1081 cP

3.2. Hydrodynamic Calculus General Considerations The process of distillation can carried out in many types of columns, the most common being the plate and the packed columns. Our project was imposed to have a rectification tower with plate trays, more specifically valve trays. Here the separation is carried out in a stage-wise manner, whereas with packed columns the process of mass transfer is continuous. The number of theoretical stages required to effect the desired separation and the corresponding rates for the liquid and vapour phases was determined by the procedures described earlier, but to translate these quantities into an actual design, the following factors have to be considered ƒthe type of tray (in our case valve tray) ƒthe vapour velocity (the major factor in determining column diameter) ƒthe 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 the liquid 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 easy to erect and maintain. 22

Unit Operations Project — Design of a rectification unit

The most important types of trays are as follows ƒbubble cap trays ƒsieve trays ƒvalve trays The latter type, which may be regarded as intermediate between the bubble cap and the sieve cap, offers advantages over both. The feature of the tray is that liftable caps act as variable orifices. A diagrammatic representation of a valve is given in Figure 7. The valves (1) are metal disks or metal stripes of up to 38 mm in diameter which are raised above the openings (3) in the tray deck (2) as vapour passes through the tray. The caps are restrained by legs or spiders which limit the vertical movement up or down. Some types are capable of forming a total liquid seal when the vapour flow is insufficient to lift the cap.

1 2

3

5

Figure 7. A valve

A cutaway section of a plate column is given in Figure 8, while a schematic representation of the most important parameters characterising a tray is represented in Figure 9. Data Inputs The computations below will be made separately (when it is necessary) for the rectifying and the stripping region of the column, at the values corresponding to the median trays. Some important data used in the calculus are gathered in the Table 7 for quick reference. Table 7. Data inputs for the hydrodynamic calculus Parameter

Symbol

Rectification

Stripping

Median tray

-

2

7

Vapour flow

QV

Liquid flow

QL

Temperature

t

324.76 kmole/h 729.00 m3/h 131.98 kmole/h 9.33 m3/h -24.2 ˚C

255.88 kmole/h 629.42 m3/h 544.77 kmole/h 44.26 m3/h 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 density Liquid density Vapour dynamic viscosity Liquid dynamic viscosity

UV UL KV KL

3

14.18 kg/m

15.31 kg/m3

504.6 kg/m3

510.3 kg/m3

8.05 E-3 cP

0.1108 cP

8.392 E-3 cP

0.1128 cP

23

Unit Operations Project — Design of a rectification unit

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

ž 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; d = valve diameter; d = hole diameter; h = height of skirt clearance of downcomer; h = weir height; H = v h cl w plate (tray) spacing; VL = volumetric flow rate of liquid B) Cross section: d) Small holes with narrow spacing; e) Large holes with wide spacing; A = cross-sectional area of downcomer; A = active cross-sectional area; D = column diameter; l = skirt d ac c cl clearance length of downcomer; lL= length of liquid path; lw= weir length

24

Unit Operations Project — Design of a rectification unit

We have chosen the following characteristics for our type of tray: Table 8. Chosen parameters Parameter Valve tray diameter Valve hole diameter Valve weight Maximum valve height Valve material Valve material density

Symbol dS d0 G1 hS -

U Cu H

Value 75 mm 65 mm 80 g 12 mm Cu 8900 kg/m3

Tray spacing Tray thickness

GS

0.3 m 5 mm

Hole area

Sr

0.15 S

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

X

§ QL · ¨¨ ¸¸ © QV ¹

Y

w 02 g ˜ de

0.25

§ UL ¨¨ © UV

§S ¨¨ © Sr

· ¸¸ ¹

0.125

2

(33)

· U V § KL · ¸¸ ¨¨ ¸¸ ¹ UL © K V ¹

0.16

(34)

and so by calculating X we can obtain Y from the diagram and further derivate the velocity as w0

Sr S

U Y ˜ g ˜ de L UV

§ KV ¨¨ © KL

· ¸¸ ¹

0.16

(35)

In the above expressions, w0 is obtained in m/s, g is the gravitational acceleration (g = 9.81 m/s), de is 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 domain for this type of trays.

Parameter X Y w0 (m/s)

Rectification 0.525 1.1 0.37

Stripping 0.8 0.4 0.22

25

Unit Operations Project — Design of a rectification unit

Column Diameter Since the vapour flow equals

S ˜ Di2 w0 4

QV

(36)

we can calculate the column diameter as 4 QV S w0

Di

(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

Sd

1 §S˜D ·  sin D ¸ ˜ S ¨ 2S © 180 ¹

(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

z

Sr S ˜ d02 4

(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 equilateral triangles. We chose ƒthe distance of the marginal valves to the column wall G1 = 50 mm ƒthe distance of the marginal valves to the weir plate G 2 = 75 mm ƒweir plate height hz = 40 mm 26

Unit Operations Project — Design of a rectification unit

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 in order to assure it doesn’t exceed a certain threshold. We use the relation

e

w m0 C ˜K n H

(40)

where C is a coefficient which for our case has the value C = 3.6 · 10-3 and K, m and n depend on the shape of the valve. For disc-like valves, the values are: K = 6.9, m = 2.7 and 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 is necessary for the pressure p1 underneath the tray to be greater than the pressure p2 above the tray. The resistance that opposes to the vapour pass equals with 'p

(41)

p1  p 2

and is composed from the resistance of the dried tray and the resistance of the liquid alone, thus

'p

(42)

'pU  'pL

The liquid component may be approximated with the relation

'pL

UL ˜ g ˜ h z  h d  0.5'

(43)

where hz is the height of the weir, hd is the height of the liquid above the weir and ' is the level difference. In its turn, hd may be deduced from the semi-empirical relation

hd

§Q · K 1 ˜ K 2 ¨¨ L ¸¸ © ld ¹

2/3

(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 is obtained as a graphical dependence between QL, ld and Di. The dry component can be calculated from the formula

'pU

w 12 [ UV 2

(45)

where w1 is the local velocity of the vapours inside the valve, 27

Unit Operations Project — Design of a rectification unit

w1

QV 3600 ˜ Sr

(46)

and [ is the hydraulic resistance coefficient of the valve, [ = 3.6. After performing the computations, we get the following results. Parameter K1 K2 ld (m) hd (mm) hz (mm) w1 (m/s) 'pL ( N/m2)

Rectification 1.02 2.84 0.707 16.2 40 1.72 288

Stripping 1.125 2.84 0.707 50.4 40 1.48 462.5

'pU ( N/m2) 'p ( N/m2)

75

60.5

363

523

Later, these values are to be multiplied with the real number of trays in order to obtain the total pressure drop through the column. We plan now to calculate the real number of trays based on the kinetics of the mass transfer. For this we need to calculate the diffusion coefficients, the partial mass transfer coefficients, the total mass transfer coefficient, and using this to draw the kinetic curve, (i.e. the real equilibrium curve) and to apply the McCabe-Thiele method for it, yielding the real number of trays. Diffusion Coefficients The values of the diffusion coefficients for the vapour and the liquid phase are dependent 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

DV

4.3 ˜ 10 7

T3 / 2





1/ 3 2 2

p ˜ V11/ 3  V

1 1  M1 M2

(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), M are the molar masses and the pressure p and the temperature T are introduced in atm and Kelvin, respectively. For the liquid phase we have used the Wilke and Chang relationship

DL

7.4 ˜ 10 14

F ˜ M 0.5 ˜ T K ˜ V00.6

(48)

28

Unit Operations Project — Design of a rectification unit

in which F denotes a parameter for association (in our case, the mixture containing only hydrocarbons, F = 0), K is the dynamic viscosity in P, M is the average molecular mass while V0 is a parameter related to the solvent (in our case V0 = 22.8). We have obtained the following results. Parameter M1 (kg/kmole) M2 (kg/kmole) M (kg/kmole) V1 (cm3/mole) V2 (cm3/mole) P (atm) KL (P) T (K) DV (m2/s) DL (m2/s)

Rectification 30 44 35.7 51.8 74.0 7.89 1.108·10-3

Stripping 30 44 35.7 51.8 74.0 7.89 1.128·10-3

248.95 8.07·10-7 1.52·10-8

274.35 9.33·10-7 1.78·10-8

Partial Mass Transfer Coefficients The equation for the mass transfer inside a phase is N

(49)

k ˜ 'c

where N is the transported mass flow (kmole/h), k is the partial mass transfer coefficient (kmole/m2h ' c) and ' c is the concentration difference. Thus, the dimensions 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 film coefficients, we have used the equation

Sh

8 ˜ 10 4 ˜ Re 0.5 ˜ Sc 0.5 ˜ *

(50)

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

8 ˜ 10 4

w ˜l˜U K hz H ˜ K U ˜ DV l

(51)

in which H denotes the fraction of holes in the suspension (we took H = 0.5) and l represents the specific dimension which we considered l = 1. Knowing the Sh criterion, it is easy to calculate the partial mass transfer coefficients since

Sh

k' V ˜ l DV

(52)

The value we obtained for k is however related to the hole area Sr and it must be connected to the total area, thus 29

Unit Operations Project — Design of a rectification unit

kV

k' V M

k' V

Sr S

(53)

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

ky

kV

1 273.15 22.41 T

(54)

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

Re

w ˜l˜U 1  H ˜ K

(55)

and kL is 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 w (m/s) U V (kg/m3)

Rectification 0.37 14.18

Stripping 0.22 15.31

UL (kg/m3) K V (P) KL (P)

504.6

510.3

1.108·10-3

1.128·10-3

ReV ReL ScV ScL DV (m2/s) DL (m2/s) ShV ShL k’V (m/h) kV (m/h) kL (m/h) ky (kmole/m2h ' y) kx (kmole/m2h ' x)

130 350 336 764 7.035 144.46 8.07·10-7 1.52·10-8 3.064·106 2.232·107 8.901·103 1335 1221.35 65.3 17 272.78

80 272 199 053 5.875 124.53 9.33·10-7 1.78·10-8 2.198·106 1.593·107 7.381·103 1107 1018.07 49.1 12 530.66

8.05·10

-5

8.39·10-5

Kinetic Curve We can now obtain, for several values of x between xD and xW, the values of the total mass transfer coefficient as

K

1 1 m  ky kx

(56)

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

m

D

>1  D  1 ˜ x@2

(57)

Now using the relation 30

Unit Operations Project — Design of a rectification unit

N

ln

y n*  y n1 y n*  y n

ln

AB BC

K ˜ Sa QV

(58)

we can obtain the coordinates of the points yn belonging to the kinetic curve. The rest of the y terms denote: yn* the equilibrium curve and yn+1 the operating lines. The computations were done in the Excel sheet presented below. xn ky (kmol/m2*h*Dy) kx (kmol/m2*h*Dx) m Ky (kmol/m2*h*Dy) V (kmol/h) Nv BC/AB y*n y n+1 AB BC AC yn

0 49.1 12530.66 4.84 48.18615 255.8797 0.121087 0.885957 0 -0.037873 0.037873 0.033554 0.004319 -0.004319

0.05 49.1 12530.66 3.406378 48.45327 255.8797 0.121758 0.885362 0.22555 0.068578 0.156973 0.138978 0.017995 0.207556

0.1 49.1 12530.66 2.526813 48.61863 255.8797 0.122174 0.884995 0.380743 0.175028 0.205715 0.182057 0.023658 0.357085

0.15 49.1 12530.66 1.948646 48.72794 255.8797 0.122448 0.884752 0.494058 0.281479 0.212579 0.18808 0.024499 0.469558

0.2 49.1 12530.66 1.548392 48.8039 255.8797 0.122639 0.884583 0.580429 0.387929 0.1925 0.170282 0.022218 0.558212

0.25 49.1 12530.66 1.2598917 48.858797 255.87966 0.1227773 0.8844606 0.6484467 0.4943798 0.1540668 0.1362661 0.0178008 0.6306459

stripping 0.3 49.1 12530.66 1.045107 48.89975 255.8797 0.12288 0.88437 0.703398 0.60083 0.102568 0.090708 0.01186 0.691538

xn ky (kmol/m2*h*Dy) kx (kmol/m2*h*Dx) m Ky (kmol/m2*h*Dy) V (kmol/h) Nv BC/AB y*n y n+1 AB BC AC yn

0.5 65.3 17272.78 0.482894 65.18101 324.7585 0.129054 0.878926 0.846944 0.770242 0.076702 0.067416 0.009287 0.837657

0.55 65.3 17272.78 0.420276 65.19641 324.7585 0.129085 0.8789 0.871188 0.790561 0.080626 0.070862 0.009764 0.861424

0.6 65.3 17272.78 0.369093 65.20901 324.7585 0.129109 0.878878 0.892477 0.810881 0.081596 0.071713 0.009883 0.882594

0.65 65.3 17272.78 0.326721 65.21944 324.7585 0.12913 0.87886 0.911321 0.831201 0.08012 0.070414 0.009706 0.901615

0.7 65.3 17272.78 0.291247 65.22818 324.7585 0.129147 0.878844 0.928118 0.85152 0.076597 0.067317 0.00928 0.918837

0.75 65.3 17272.78 0.2612526 65.235569 324.75847 0.1291621 0.8788315 0.943184 0.87184 0.071344 0.0626994 0.0086446 0.9345393

0.8 65.3 17272.78 0.235664 65.24187 324.7585 0.129175 0.878821 0.956774 0.89216 0.064614 0.056784 0.00783 0.948944

feed 0.35 49.1 12530.66 0.880907 48.9311 255.8797 0.122959 0.8843 0.748719 0.707281 0.041438 0.036644 0.004794 0.743925

rectifying 0.4 65.3 17272.78 0.658751 65.13778 324.7585 0.128968 0.879002 0.786737 0.729603 0.057134 0.050221 0.006913 0.779823

0.45 65.3 17272.78 0.560624 65.16189 324.7585 0.129016 0.87896 0.819085 0.749922 0.069163 0.060791 0.008371 0.810713

0.85 65.3 17272.78 0.213658 65.2473 324.7585 0.129185 0.878811 0.969095 0.912479 0.056615 0.049754 0.006861 0.962233

0.9 65.3 17272.78 0.194597 65.252 324.7585 0.129195 0.878803 0.980316 0.932799 0.047517 0.041758 0.005759 0.974557

0.95 65.3 17272.78 0.177977 65.25609 324.7585 0.129203 0.878796 0.990578 0.953118 0.03746 0.03292 0.00454 0.986038

1 65.3 17272.78 0.163399 65.25969 324.7585 0.12921 0.87879 1 0.973438 0.026562 0.023342 0.00322 0.99678

The kinetic curve obtained is presented in Figure 10. Applying again the McCabeThiele method, this time between the operating lines and the newly drawn kinetic curve, 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 the pressure drop per tray and the real number of trays in the specified section of the column (rectifying or stripping). We obtain

'p = 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 are required to calculate the height of the column and to draw a scale representation of it.

31

Unit Operations Project — Design of a rectification unit

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

1

0.9

0.8

0.7

y

0.6

0.5

0.4

0.3

0.2

0.1

0 0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

x

The height of the column The total height of the column is computed by the approximate formula Ht # Ha + HV + HB + HS

(59)

where HV is the distance from the first tray to the top of the column, HB is the distance from the last tray to the bottom of the column, HS is the distance from the bottom of the column to the ground, and Ha is the height of the active section of the column Ha = (NRT-1)·H + NRT· G S

(60)

in which H is the distance between the trays and G S is the tray thickness. Note that since the reboiler works at equilibrium, the trays to be actually considered in equation 32

Unit Operations Project — Design of a rectification unit

(60) are one less (the reboiler acts as a tray but is not computed in the total height of the column). We have considered H = 0.3 m

G 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 of Bucharest, 2004. 2. A. Woinaroschy — Unit Operations in Chemical Engineering, Politehnica University Press, Bucharest, 1994. 3. O. Floarea, G. Jinescu, C. Balaban, P. Vasilescu, R. Dima — Operatii si Utilaje in Industria Chimica. Probleme pentru subingineri, Editura Didactica si Pedagogica, Bucuresti, 1980. 4. * * * — Ingineria Proceselor Fizice si Chimice. Ghid de proiect, Politehnica University Press, Bucharest.

33

Appendixes

Case Name:

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

Unit Set:

SI

Date/Time:

Thu Jun 03 19:29:47 2004

TEAM LND Calgary, Alberta CANADA

Pure Component: Ethane Identification Family / Class

Chemical Formula

Hydrocarbon

ID Number

C2H6

Group Name

CAS Number

2

---

UNIFAC Structure (CH3)2

User ID Tags Tag Number

Tag Text

Critical/Base Properties Base Properties

Critical Properties

Molecular Weight

30.07

Normal Boiling Pt Std Liq Density

(C)

-88.60

(kg/m3)

355.7

Temperature

(C)

Pressure Volume

32.28

(kPa)

4884

(m3/kgmole)

0.1480

Acentricity

9.860e-002

Temperature Dependent Properties Vapour Enthalpy

Vapour Pressure

Minimum Temperature

(C)

-270.0

-140.1

Maximum Temperature

(C)

5000

32.25

Coefficient Name

IdealH Coefficient

Antoine Coefficient

Gibbs Free Energy 25.00 426.9 Gibbs Free Coefficient

a

-1.768

44.01

b

1.143

-2569

-8.579e+004 168.6

c

-3.236e-004

0.0000

2.685e-002

d

4.243e-006

-4.976

0.0000

e

-3.393e-009

1.464e-005

0.0000

f

8.821e-013

2.000

---

g

1.000

0.0000

---

h

0.0000

0.0000

---

i

0.0000

0.0000

---

j

0.0000

0.0000

---

Additional Point Properties Thermodynamic and Physical Properties Dipole Moment Radius of Gyration

1.826

COSTALD (SRK) Acentricity COSTALD Volume

Property Package Molecular Properties 0.0000

9.830e-002 (m3/kgmole)

0.1458

PRSV - Kappa

0.0000

GS/CS - Solubility Parameter

7.231e-002

GS/CS - Molar Volume

Viscosity Coefficient B

4.698e-002

GS/CS - Acentricity

Cavett Heat of Vap Coeff A

0.2833

UNIQUAC - R

Cavett Heat of Vap Coeff B

---

UNIQUAC - Q

Heat of Formation (25C)

(kJ/kgmole)

-8.474e+004

Wilson Molar Volume

Heat of Combustion (25C)

(kJ/kgmole)

-1.429e+006

CN Solubility

Enthalpy Basis Offset

(kJ/kgmole)

-9.670e+004

CN Molar Volume

Licensed to: TEAM LND

2.198

ZJ EOS Parameter

Viscosity Coefficient A

Hyprotech Ltd.

1.343e-002

KD Group Parameter

HYSYS v3.0.1 (Build 4602)

6.050 (m3/kgmole)

6.700e-002 0.1064 1.802 1.696

(m3/kgmole)

8.454e-002

(m3/kgmole)

9.798e-003

4.287

Page 1 of 1

Case Name:

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

Unit Set:

SI

Date/Time:

Thu Jun 03 19:28:57 2004

TEAM LND Calgary, Alberta CANADA

Pure Component: Propane Identification Family / Class

Chemical Formula

Hydrocarbon

ID Number

C3H8

Group Name

CAS Number

3

---

UNIFAC Structure (CH3)2 CH2

User ID Tags Tag Number

Tag Text

Critical/Base Properties Base Properties

Critical Properties

Molecular Weight

44.10

Normal Boiling Pt Std Liq Density

(C)

-42.10

(kg/m3)

506.7

Temperature

(C)

Pressure Volume

96.75

(kPa)

4257

(m3/kgmole)

0.2000

Acentricity

0.1524

Temperature Dependent Properties Vapour Enthalpy

Vapour Pressure

Minimum Temperature

(C)

-270.0

-128.1

Maximum Temperature

(C)

5000

96.65

Coefficient Name

IdealH Coefficient

Antoine Coefficient

Gibbs Free Energy 25.00 426.9 Gibbs Free Coefficient

a

39.49

52.38

b

0.3950

-3491

-1.055e+005 264.8

c

2.114e-003

0.0000

3.250e-002

d

3.965e-007

-6.109

0.0000

e

-6.672e-010

1.119e-005

0.0000

f

1.679e-013

2.000

---

g

1.000

0.0000

---

h

0.0000

0.0000

---

i

0.0000

0.0000

---

j

0.0000

0.0000

---

Additional Point Properties Thermodynamic and Physical Properties Dipole Moment Radius of Gyration

2.431

COSTALD (SRK) Acentricity COSTALD Volume

Property Package Molecular Properties 0.0000

(m3/kgmole)

PRSV - Kappa

0.1532

ZJ EOS Parameter

0.2001

GS/CS - Solubility Parameter

Viscosity Coefficient A

7.112e-002

Viscosity Coefficient B

-6.538e-002

GS/CS - Molar Volume

0.2783

UNIQUAC - R

Cavett Heat of Vap Coeff B

---

UNIQUAC - Q

Heat of Formation (25C)

(kJ/kgmole)

-1.039e+005

Wilson Molar Volume

Heat of Combustion (25C)

(kJ/kgmole)

-2.045e+006

CN Solubility

Enthalpy Basis Offset

(kJ/kgmole)

-1.194e+005

CN Molar Volume

Licensed to: TEAM LND

3.007 0.0000 6.400 (m3/kgmole)

GS/CS - Acentricity

Cavett Heat of Vap Coeff A

Hyprotech Ltd.

3.163e-002

KD Group Parameter

HYSYS v3.0.1 (Build 4602)

8.400e-002 0.1538 2.477 2.236

(m3/kgmole)

8.703e-002

(m3/kgmole)

1.072e-002

5.999

Page 1 of 1

Case Name:

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

Unit Set:

SI

Date/Time:

Thu Jun 03 19:30:08 2004

TEAM LND Calgary, Alberta CANADA

Pure Component: n-Butane Identification Family / Class

Chemical Formula

Hydrocarbon

ID Number

C4H10

Group Name

CAS Number

5

---

UNIFAC Structure (CH3)2 (CH2)2

User ID Tags Tag Number

Tag Text

Critical/Base Properties Base Properties

Critical Properties

Molecular Weight

58.12

Normal Boiling Pt Std Liq Density

(C)

-0.5020

(kg/m3)

583.2

Temperature

(C)

Pressure Volume

152.0

(kPa)

3797

(m3/kgmole)

0.2550

Acentricity

0.2010

Temperature Dependent Properties Vapour Enthalpy

Vapour Pressure

Minimum Temperature

(C)

-270.0

-103.1

Maximum Temperature

(C)

5000

152.0

Coefficient Name

IdealH Coefficient

Antoine Coefficient

Gibbs Free Energy 25.00 426.9 Gibbs Free Coefficient

a

67.72

66.94

b

8.541e-003

-4604

-1.284e+005 360.5

c

3.277e-003

0.0000

3.826e-002

d

-1.110e-006

-8.255

0.0000

e

1.766e-010

1.157e-005

0.0000

f

-6.399e-015

2.000

---

g

1.000

0.0000

---

h

0.0000

0.0000

---

i

0.0000

0.0000

---

j

0.0000

0.0000

---

Additional Point Properties Thermodynamic and Physical Properties Dipole Moment Radius of Gyration

2.886

COSTALD (SRK) Acentricity COSTALD Volume

Property Package Molecular Properties 0.0000

(m3/kgmole)

PRSV - Kappa

0.2008

ZJ EOS Parameter

0.2544

GS/CS - Solubility Parameter

Viscosity Coefficient A

0.1001

GS/CS - Molar Volume

Viscosity Coefficient B

-5.969e-002 0.2747

UNIQUAC - R

Cavett Heat of Vap Coeff B

---

UNIQUAC - Q

Heat of Formation (25C)

(kJ/kgmole)

-1.262e+005

Wilson Molar Volume

Heat of Combustion (25C)

(kJ/kgmole)

-2.660e+006

CN Solubility

Enthalpy Basis Offset

(kJ/kgmole)

-1.456e+005

CN Molar Volume

Licensed to: TEAM LND

4.027 0.0000 6.730 (m3/kgmole)

GS/CS - Acentricity

Cavett Heat of Vap Coeff A

Hyprotech Ltd.

3.951e-002

KD Group Parameter

HYSYS v3.0.1 (Build 4602)

0.1014 0.1953 3.151 2.776

(m3/kgmole)

9.966e-002

(m3/kgmole)

1.277e-002

6.713

Page 1 of 1

Case Name:

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

Unit Set:

SI

Date/Time:

Wed Jun 02 20:26:56 2004

TEAM LND Calgary, Alberta CANADA

Distillation: First Column @Main CONNECTIONS Inlet Stream

Stage

STREAM NAME Qreb1 F1

FROM UNIT OPERATION

Reboiler 6__Main TS Outlet Stream

Stage

STREAM NAME Qcond1 D1 W1

TO UNIT OPERATION

Condenser Condenser Reboiler

Valve

VLV-100

MONITOR Specifications Summary Specified Value

Wt. Error

Wt. 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

479.8 kgmole/h

481.7 kgmole/h

3.901e-003

1.000e-002

1.000 kgmole/h

Off

On

Off

6.000

2.841

-0.5265

1.000e-002

1.000e-002

Off

On

Off

Distillate Rate Reflux Ratio

Current Value

Abs. Tol.

PROPERTIES Properties : F1 Overall Vapour/Phase Fraction Temperature: Pressure: Molar Flow Mass Flow Std Ideal Liq Vol Flow Molar Enthalpy Mass Enthalpy

Vapour Phase

Liquid Phase

0.5500

0.5500

0.4500

(C)

59.67

59.67

59.67

(kPa)

1400

1400

1400

(kgmole/h)

950.0

522.5

427.5 2.230e+004

(kg/h)

4.589e+004

2.359e+004

(m3/h)

88.21

47.52

40.69

(kJ/kgmole)

-1.171e+005

-1.053e+005

-1.314e+005

(kJ/kg)

-2423

-2333

-2519

Molar Entropy

(kJ/kgmole-C)

123.3

152.0

88.33

Mass Entropy

(kJ/kg-C)

2.553

3.367

1.693

(kJ/h)

-1.112e+008

-5.503e+007

-5.619e+007

Molar Density

(kgmole/m3)

1.093

0.6341

9.500

Mass Density

(kg/m3)

52.81

28.62

495.7

Std Ideal Liq Mass Density

(kg/m3)

520.3

496.4

548.2

Liq Mass Density @Std Cond (kg/m3)

537.4

517.9

557.9

Molar Heat Capacity

(kJ/kgmole-C)

122.4

95.08

155.8

Mass Heat Capacity

(kJ/kg-C)

2.534

2.106

2.986

Thermal Conductivity

(W/m-K)

---

2.302e-002

7.792e-002

(cP)

---

9.945e-003

0.1012

(dyne/cm)

---

---

6.136

48.31

45.14

52.17

---

0.7979

5.326e-002

Heat Flow

Viscosity Surface Tension Molecular Weight Z Factor

Properties : D1 Overall Vapour/Phase Fraction

Hyprotech Ltd. Licensed to: TEAM LND

0.0000

Vapour Phase 0.0000

HYSYS v3.0.1 (Build 4602)

Liquid Phase 1.0000

Page 1 of 3

Case Name:

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

Unit Set:

SI

Date/Time:

Wed Jun 02 20:26:56 2004

TEAM LND Calgary, Alberta CANADA

Distillation: First Column @Main (continued) Properties : D1 Overall Temperature: Pressure: Molar Flow Mass Flow Std Ideal Liq Vol Flow Molar Enthalpy Mass Enthalpy

Vapour Phase

Liquid Phase

(C)

9.259

9.259

(kPa)

1400

1400

9.259 1400

(kgmole/h)

481.7

0.0000

481.7 1.871e+004

(kg/h)

1.871e+004

0.0000

(m3/h)

41.57

0.0000

41.57

(kJ/kgmole)

-1.125e+005

-9.317e+004

-1.125e+005

(kJ/kg)

-2898

-2693

-2898

Molar Entropy

(kJ/kgmole-C)

106.5

161.8

106.5

Mass Entropy

(kJ/kg-C)

2.743

4.676

2.743

(kJ/h)

-5.421e+007

0.0000

-5.421e+007

Molar Density

(kgmole/m3)

12.31

0.7443

12.31

Mass Density

(kg/m3)

478.1

25.75

478.1

Std Ideal Liq Mass Density

(kg/m3)

450.1

405.2

450.1

Liq Mass Density @Std Cond (kg/m3)

468.3

425.5

468.3

Molar Heat Capacity

(kJ/kgmole-C)

115.9

68.11

115.9

Mass Heat Capacity

(kJ/kg-C)

2.983

1.969

2.983

Thermal Conductivity

(W/m-K)

9.416e-002

2.020e-002

9.416e-002

(cP)

9.235e-002

9.188e-003

9.235e-002

(dyne/cm)

6.250

---

6.250

38.84

34.60

38.84

4.843e-002

0.8011

4.843e-002

Heat Flow

Viscosity Surface Tension Molecular Weight Z Factor

Properties : W1 Overall Vapour/Phase Fraction Temperature: Pressure: Molar Flow Mass Flow Std Ideal Liq Vol Flow Molar Enthalpy Mass Enthalpy

Vapour Phase

Liquid Phase

0.0000

0.0000

1.0000

(C)

94.61

94.61

94.61

(kPa)

1400

1400

1400

(kgmole/h)

468.3

0.0000

468.3 2.718e+004

(kg/h)

2.718e+004

0.0000

(m3/h)

46.64

0.0000

46.64

(kJ/kgmole)

-1.365e+005

-1.208e+005

-1.365e+005

(kJ/kg)

-2351

-2084

-2351

Molar Entropy

(kJ/kgmole-C)

82.33

125.0

82.33

Mass Entropy

(kJ/kg-C)

1.419

2.157

1.419

(kJ/h)

-6.391e+007

0.0000

-6.391e+007

Molar Density

(kgmole/m3)

8.196

0.6164

8.196

Mass Density

(kg/m3)

475.7

35.72

475.7

Std Ideal Liq Mass Density

(kg/m3)

582.8

582.4

582.8

Liq Mass Density @Std Cond (kg/m3)

584.2

583.8

584.2

Molar Heat Capacity

(kJ/kgmole-C)

190.3

133.8

190.3

Mass Heat Capacity

(kJ/kg-C)

3.278

2.309

3.278

Thermal Conductivity

(W/m-K)

6.687e-002

2.432e-002

6.687e-002

(cP)

9.060e-002

9.996e-003

9.060e-002

(dyne/cm)

4.503

---

4.503

58.04

57.95

58.04

5.586e-002

0.7428

5.586e-002

Heat Flow

Viscosity Surface Tension Molecular Weight Z Factor

Hyprotech Ltd. Licensed to: TEAM LND

HYSYS v3.0.1 (Build 4602)

Page 2 of 3

Case Name:

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

Unit Set:

SI

Date/Time:

Wed Jun 02 20:26:56 2004

TEAM LND Calgary, Alberta CANADA

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 Temperature (C)

Pressure (kPa)

Net Liq (kgmole/h)

Net Vap (kgmole/h)

Net Feed (kgmole/h)

Net Draws (kgmole/h)

Condenser

9.259

1400

1368

---

---

481.7

1__Main TS

26.84

1400

1329

1850

---

---

2__Main TS

37.52

1400

1289

1811

---

---

3__Main TS

45.61

1400

1234

1770

---

---

4__Main TS

53.29

1400

1188

1716

---

---

5__Main TS

60.15

1400

1163

1670

---

---

6__Main TS

65.34

1400

1599

1645

950.0

---

7__Main TS

72.92

1400

1636

1130

---

---

8__Main TS

78.29

1400

1659

1168

---

---

9__Main TS

82.47

1400

1675

1191

---

---

10__Main TS

85.83

1400

1690

1207

---

---

11__Main TS

88.47

1400

1702

1221

---

---

12__Main TS

90.48

1400

1713

1234

---

---

13__Main TS

91.96

1400

1721

1244

---

---

14__Main TS

93.02

1400

1727

1253

---

---

15__Main TS

93.76

1400

1732

1259

---

---

16__Main TS

94.27

1400

1735

1263

---

---

Reboiler

94.61

1400

---

1267

---

468.3

Column Profiles Energy Temperature (C)

Liquid Enthalpy (kJ/kgmole)

Vapour Enthalpy (kJ/kgmole)

Heat Loss (kJ/h)

Condenser

9.259

-1.125e+005

-9.317e+004

---

1__Main TS

26.84

-1.174e+005

-9.813e+004

---

2__Main TS

37.52

-1.207e+005

-1.014e+005

---

3__Main TS

45.61

-1.239e+005

-1.034e+005

---

4__Main TS

53.29

-1.271e+005

-1.052e+005

---

5__Main TS

60.15

-1.296e+005

-1.069e+005

---

6__Main TS

65.34

-1.313e+005

-1.084e+005

---

7__Main TS

72.92

-1.324e+005

-1.117e+005

---

8__Main TS

78.29

-1.334e+005

-1.139e+005

---

9__Main TS

82.47

-1.342e+005

-1.156e+005

---

10__Main TS

85.83

-1.349e+005

-1.170e+005

---

11__Main TS

88.47

-1.354e+005

-1.181e+005

---

12__Main TS

90.48

-1.358e+005

-1.189e+005

---

13__Main TS

91.96

-1.360e+005

-1.196e+005

---

14__Main TS

93.02

-1.362e+005

-1.201e+005

---

15__Main TS

93.76

-1.363e+005

-1.204e+005

---

16__Main TS

94.27

-1.364e+005

-1.206e+005

---

Reboiler

94.61

-1.365e+005

-1.208e+005

---

Hyprotech Ltd. Licensed to: TEAM LND

HYSYS v3.0.1 (Build 4602)

Page 3 of 3

Case Name:

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

Unit Set:

SI

Date/Time:

Wed Jun 02 20:17:26 2004

TEAM LND Calgary, Alberta CANADA

HYSYS Column Profiles Specsheet: Column Temperature Profile

Temperature vs. Tray Number 100 90.0

Temperature (C)

80.0 70.0 60.0 50.0 40.0 30.0 20.0 10.0 0.000 0

2

4

6

8

10

12

14

16

18

Column Temperature Profile Column Stage

Temperature (C)

Hyprotech Ltd. Licensed to: TEAM LND

Condenser

9.259

1__Main TS

26.84

2__Main TS

37.52

3__Main TS

45.61

4__Main TS

53.29

5__Main TS

60.15

6__Main TS

65.34

7__Main TS

72.92

8__Main TS

78.29

9__Main TS

82.47

10__Main TS

85.83

11__Main TS

88.47

12__Main TS

90.48

13__Main TS

91.96

14__Main TS

93.02

15__Main TS

93.76

16__Main TS

94.27

Reboiler

94.61

HYSYS v3.0.1 (Build 4602)

Page 1 of 1

Case Name:

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

Unit Set:

SI

Date/Time:

Wed Jun 02 20:21:21 2004

TEAM LND Calgary, Alberta CANADA

HYSYS Column Profiles Specsheet:

Column Properties Profile Options Selected Mass basis is selected Stage

Surface Tension

Mole Weight (Vap)

Density (Vap)

Viscosity (Vap)

Therm Cond (Vap)

Heat Cap (Vap)

Mole Weight (Lt Liq)

Condenser

6.250 dyne/cm

34.60

25.75 kg/m3

9.188e-003 cP

2.020e-002 W/m-K

1.969 kJ/kg-C

38.84

1__Main TS

5.921 dyne/cm

38.84

27.53 kg/m3

9.407e-003 cP

2.094e-002 W/m-K

2.016 kJ/kg-C

42.60

2__Main TS

5.777 dyne/cm

41.60

28.72 kg/m3

9.514e-003 cP

2.140e-002 W/m-K

2.051 kJ/kg-C

45.08

3__Main TS

5.760 dyne/cm

43.38

29.22 kg/m3

9.622e-003 cP

2.185e-002 W/m-K

2.075 kJ/kg-C

47.33

4__Main TS

5.757 dyne/cm

44.95

29.53 kg/m3

9.737e-003 cP

2.234e-002 W/m-K

2.099 kJ/kg-C

49.52

5__Main TS

5.713 dyne/cm

46.44

29.90 kg/m3

9.836e-003 cP

2.277e-002 W/m-K

2.122 kJ/kg-C

51.35

6__Main TS

5.646 dyne/cm

47.69

30.28 kg/m3

9.902e-003 cP

2.308e-002 W/m-K

2.142 kJ/kg-C

52.63

7__Main TS

5.271 dyne/cm

50.39

31.76 kg/m3

9.915e-003 cP

2.335e-002 W/m-K

2.184 kJ/kg-C

53.90

8__Main TS

5.054 dyne/cm

52.24

32.72 kg/m3

9.933e-003 cP

2.358e-002 W/m-K

2.213 kJ/kg-C

54.89

9__Main TS

4.910 dyne/cm

53.65

33.44 kg/m3

9.951e-003 cP

2.377e-002 W/m-K

2.236 kJ/kg-C

55.70

10__Main TS

4.801 dyne/cm

54.79

34.03 kg/m3

9.966e-003 cP

2.392e-002 W/m-K

2.255 kJ/kg-C

56.36

11__Main TS

4.715 dyne/cm

55.71

34.50 kg/m3

9.977e-003 cP

2.405e-002 W/m-K

2.271 kJ/kg-C

56.87

12__Main TS

4.649 dyne/cm

56.43

34.88 kg/m3

9.985e-003 cP

2.414e-002 W/m-K

2.283 kJ/kg-C

57.26

13__Main TS

4.598 dyne/cm

56.96

35.17 kg/m3

9.989e-003 cP

2.421e-002 W/m-K

2.292 kJ/kg-C

57.54

14__Main TS

4.561 dyne/cm

57.35

35.39 kg/m3

9.992e-003 cP

2.425e-002 W/m-K

2.299 kJ/kg-C

57.74

15__Main TS

4.534 dyne/cm

57.63

35.54 kg/m3

9.994e-003 cP

2.429e-002 W/m-K

2.303 kJ/kg-C

57.88

16__Main TS

4.515 dyne/cm

57.82

35.65 kg/m3

9.996e-003 cP

2.431e-002 W/m-K

2.307 kJ/kg-C

57.97

Reboiler

4.503 dyne/cm

57.95

35.72 kg/m3

9.996e-003 cP

2.432e-002 W/m-K

2.309 kJ/kg-C

58.04

Stage

Surface Tension

Density (Lt Liq)

Viscosity (Lt Liq)

Therm Cond (Lt Liq)

Heat Cap (Lt Liq)

Condenser

6.250 dyne/cm

478.1 kg/m3

9.235e-002 cP

9.416e-002 W/m-K

2.983 kJ/kg-C

1__Main TS

5.921 dyne/cm

479.3 kg/m3

9.180e-002 cP

8.837e-002 W/m-K

3.020 kJ/kg-C

2__Main TS

5.777 dyne/cm

480.6 kg/m3

9.165e-002 cP

8.523e-002 W/m-K

3.039 kJ/kg-C

3__Main TS

5.760 dyne/cm

484.2 kg/m3

9.389e-002 cP

8.251e-002 W/m-K

3.033 kJ/kg-C

4__Main TS

5.757 dyne/cm

487.5 kg/m3

9.592e-002 cP

7.982e-002 W/m-K

3.029 kJ/kg-C

5__Main TS

5.713 dyne/cm

489.0 kg/m3

9.717e-002 cP

7.759e-002 W/m-K

3.039 kJ/kg-C

6__Main TS

5.646 dyne/cm

489.2 kg/m3

9.723e-002 cP

7.602e-002 W/m-K

3.054 kJ/kg-C

7__Main TS

5.271 dyne/cm

485.0 kg/m3

9.479e-002 cP

7.345e-002 W/m-K

3.114 kJ/kg-C

8__Main TS

5.054 dyne/cm

482.6 kg/m3

9.349e-002 cP

7.168e-002 W/m-K

3.153 kJ/kg-C

9__Main TS

4.910 dyne/cm

481.0 kg/m3

9.267e-002 cP

7.039e-002 W/m-K

3.183 kJ/kg-C

10__Main TS

4.801 dyne/cm

479.6 kg/m3

9.207e-002 cP

6.939e-002 W/m-K

3.207 kJ/kg-C

11__Main TS

4.715 dyne/cm

478.5 kg/m3

9.160e-002 cP

6.862e-002 W/m-K

3.227 kJ/kg-C

12__Main TS

4.649 dyne/cm

477.6 kg/m3

9.126e-002 cP

6.805e-002 W/m-K

3.243 kJ/kg-C

13__Main TS

4.598 dyne/cm

476.9 kg/m3

9.101e-002 cP

6.763e-002 W/m-K

3.255 kJ/kg-C

14__Main TS

4.561 dyne/cm

476.4 kg/m3

9.083e-002 cP

6.732e-002 W/m-K

3.264 kJ/kg-C

15__Main TS

4.534 dyne/cm

476.1 kg/m3

9.072e-002 cP

6.711e-002 W/m-K

3.271 kJ/kg-C

16__Main TS

4.515 dyne/cm

475.8 kg/m3

9.064e-002 cP

6.697e-002 W/m-K

3.275 kJ/kg-C

Reboiler

4.503 dyne/cm

475.7 kg/m3

9.060e-002 cP

6.687e-002 W/m-K

3.278 kJ/kg-C

Hyprotech Ltd. Licensed to: TEAM LND

HYSYS v3.0.1 (Build 4602)

Page 1 of 1

Case Name:

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

Unit Set:

SI

Date/Time:

Wed Jun 02 20:20:09 2004

TEAM LND Calgary, Alberta CANADA

HYSYS Column Profiles Specsheet:

Composition vs. Tray Number 1.00 Ethane (Vap) 0.900

Propane (Vap) n-Butane (Vap)

0.800

Ethane (Light) Propane (Light)

0.700

Mole Fraction

n-Butane (Light) 0.600 0.500 0.400 0.300 0.200 1.00e-001 0.000 0

2

4

6

8

10

12

14

16

18

Column Composition Profile Options Selected Fraction is selected as the composition basis

Net is selected as flow basis

Molar basis is selected Stage

Ethane (Light Liq)

Propane (Light Liq)

n-Butane (Light Liq)

Condenser

Ethane (Vap) 0.6808

0.3158

0.0034

0.3945

0.5857

0.0198

1__Main TS

0.3945

0.5857

0.0198

0.1779

0.7506

0.0715

2__Main TS

0.2355

0.7068

0.0577

0.0925

0.7451

0.1623

3__Main TS

0.1747

0.7018

0.1236

0.0619

0.6460

0.2922

4__Main TS

0.1552

0.6291

0.2157

0.0500

0.5131

0.4369

5__Main TS

0.1494

0.5340

0.3166

0.0446

0.3936

0.5618

6__Main TS

0.1471

0.4499

0.4031

0.0417

0.3084

0.6500

7__Main TS

0.0589

0.4336

0.5075

0.0158

0.2699

0.7144

8__Main TS

0.0221

0.3756

0.6023

0.0057

0.2193

0.7750

9__Main TS

0.0079

0.3032

0.6889

0.0020

0.1689

0.8292

10__Main TS

0.0027

0.2320

0.7653

0.0007

0.1246

0.8747

11__Main TS

0.0009

0.1701

0.8290

0.0002

0.0889

0.9109

12__Main TS

0.0003

0.1203

0.8794

0.0001

0.0616

0.9383

13__Main TS

0.0001

0.0825

0.9174

0.0000

0.0416

0.9584

14__Main TS

0.0000

0.0549

0.9451

0.0000

0.0274

0.9726

15__Main TS

0.0000

0.0353

0.9646

0.0000

0.0175

0.9825

16__Main TS

0.0000

0.0218

0.9782

0.0000

0.0107

0.9893

Reboiler

0.0000

0.0124

0.9876

0.0000

0.0061

Hyprotech Ltd. Licensed to: TEAM LND

Propane (Vap)

n-Butane (Vap)

HYSYS v3.0.1 (Build 4602)

0.9939

Page 1 of 1

Case Name:

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

Unit Set:

SI

Date/Time:

Wed Jun 02 20:46:02 2004

TEAM LND Calgary, Alberta CANADA

Distillation: Second Column (binary) @Main CONNECTIONS Inlet Stream

Stage

STREAM NAME Qcond22 F22

FROM UNIT OPERATION

Reboiler 4__Main TS Outlet Stream

Stage

STREAM NAME Qreb22 D22 W22

TO UNIT OPERATION

Condenser Condenser Reboiler MONITOR Specifications Summary Specified Value

Active

Estimate

Used

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

Reflux Ratio

Current Value

Wt. Error

Wt. Tol.

Abs. Tol.

PROPERTIES Properties : F22 Overall Vapour/Phase Fraction

Vapour Phase

Liquid Phase

0.1433

0.1433

0.8567

(C)

-10.76

-10.76

-10.76

(kPa)

800.0

800.0

800.0

Molar Flow

(kgmole/h)

481.7

69.02

412.7

Mass Flow

(kg/h)

1.852e+004

2374

1.615e+004

Temperature: Pressure:

Std Ideal Liq Vol Flow Molar Enthalpy Mass Enthalpy

(m3/h)

41.44

5.888

35.55

(kJ/kgmole)

-1.121e+005

-9.348e+004

-1.152e+005

(kJ/kg)

-2915

-2717

-2944

Molar Entropy

(kJ/kgmole-C)

106.4

163.7

96.82

Mass Entropy

(kJ/kg-C)

2.767

4.760

2.474

(kJ/h)

-5.399e+007

-6.451e+006

-4.754e+007

Molar Density

(kgmole/m3)

2.469

0.4221

13.07

Mass Density

(kg/m3)

94.93

14.52

511.3

Std Ideal Liq Mass Density

(kg/m3)

447.0

403.3

454.2

Liq Mass Density @Std Cond (kg/m3)

464.7

423.2

470.6

Molar Heat Capacity

(kJ/kgmole-C)

98.59

60.26

105.0

Mass Heat Capacity

(kJ/kg-C)

2.564

1.752

2.683

Thermal Conductivity

(W/m-K)

---

1.741e-002

0.1086

(cP)

---

8.277e-003

0.1146

(dyne/cm)

---

---

8.901

38.45

34.40

39.13

---

0.8688

2.806e-002

Heat Flow

Viscosity Surface Tension Molecular Weight Z Factor

Hyprotech Ltd. Licensed to: TEAM LND

HYSYS v3.0.1 (Build 4602)

Page 1 of 3

Case Name:

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

Unit Set:

SI

Date/Time:

Wed Jun 02 20:46:02 2004

TEAM LND Calgary, Alberta CANADA

Distillation: Second Column (binary) @Main (c Properties : D22 Overall Vapour/Phase Fraction

Vapour Phase

Liquid Phase

0.0000

0.0000

1.0000

(C)

-37.85

-37.85

-37.85

(kPa)

800.0

800.0

800.0

Molar Flow

(kgmole/h)

192.8

0.0000

192.8

Mass Flow

(kg/h)

5918

0.0000

5918

(m3/h)

16.32

0.0000

16.32

Molar Enthalpy

(kJ/kgmole)

-1.021e+005

-8.875e+004

-1.021e+005

Mass Enthalpy

(kJ/kg)

-3326

-2940

-3326

Molar Entropy

(kJ/kgmole-C)

111.1

164.6

111.1

Mass Entropy

(kJ/kg-C)

3.619

5.451

3.619

(kJ/h)

-1.968e+007

0.0000

-1.968e+007

Molar Density

(kgmole/m3)

15.66

0.4764

15.66

Mass Density

(kg/m3)

480.6

14.38

480.6

Std Ideal Liq Mass Density

(kg/m3)

362.6

357.0

362.6

Liq Mass Density @Std Cond (kg/m3)

370.8

361.7

370.8

Molar Heat Capacity

(kJ/kgmole-C)

88.11

52.02

88.11

Mass Heat Capacity

(kJ/kg-C)

2.870

1.723

2.870

Thermal Conductivity

(W/m-K)

0.1187

1.598e-002

0.1187

(cP)

9.783e-002

7.732e-003

9.783e-002

(dyne/cm)

8.480

---

8.480

30.70

30.19

30.70

2.612e-002

0.8584

2.612e-002

Temperature: Pressure:

Std Ideal Liq Vol Flow

Heat Flow

Viscosity Surface Tension Molecular Weight Z Factor

Properties : W22 Overall Vapour/Phase Fraction

Vapour Phase

Liquid Phase

0.0000

0.0000

1.0000

(C)

14.80

14.80

14.80

(kPa)

800.0

800.0

800.0

Molar Flow

(kgmole/h)

288.9

0.0000

288.9

Mass Flow

(kg/h)

1.260e+004

0.0000

1.260e+004

Temperature: Pressure:

Std Ideal Liq Vol Flow Molar Enthalpy Mass Enthalpy

(m3/h)

25.12

0.0000

25.12

(kJ/kgmole)

-1.205e+005

-1.037e+005

-1.205e+005

(kJ/kg)

-2761

-2430

-2761

Molar Entropy

(kJ/kgmole-C)

88.85

146.1

88.85

Mass Entropy

(kJ/kg-C)

2.037

3.425

2.037

(kJ/h)

-3.480e+007

0.0000

-3.480e+007

Molar Density

(kgmole/m3)

11.57

0.3940

11.57

Mass Density

(kg/m3)

504.9

16.81

504.9

Std Ideal Liq Mass Density

(kg/m3)

501.7

491.6

501.7

Liq Mass Density @Std Cond (kg/m3)

504.6

498.0

504.6

Molar Heat Capacity

(kJ/kgmole-C)

121.6

77.61

121.6

Mass Heat Capacity

(kJ/kg-C)

2.788

1.819

2.788

Thermal Conductivity

(W/m-K)

9.994e-002

1.799e-002

9.994e-002

(cP)

0.1081

8.390e-003

0.1081

(dyne/cm)

7.895

---

7.895

43.63

42.66

43.63

2.887e-002

0.8482

2.887e-002

Heat Flow

Viscosity Surface Tension Molecular Weight Z Factor

Hyprotech Ltd. Licensed to: TEAM LND

HYSYS v3.0.1 (Build 4602)

Page 2 of 3

Case Name:

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

Unit Set:

SI

Date/Time:

Wed Jun 02 20:46:02 2004

TEAM LND Calgary, Alberta CANADA

Distillation: Second Column (binary) @Main (c COLUMN PROFILES Reflux Ratio:

1.432

Reboil Ratio:

1.209

The Flows Option is Selected

Flow Basis:

Molar

Column Profiles Flows Temperature (C)

Pressure (kPa)

Net Liq (kgmole/h)

Net Vap (kgmole/h)

Net Feed (kgmole/h)

Net Draws (kgmole/h)

Condenser

-37.85

800.0

276.1

---

---

192.8

1__Main TS

-32.80

800.0

251.8

468.9

---

---

2__Main TS

-23.89

800.0

230.4

444.6

---

---

3__Main TS

-15.03

800.0

220.8

423.2

---

---

4__Main TS

-9.595

800.0

630.0

413.6

481.7

---

5__Main TS

-5.233

800.0

627.4

341.0

---

---

6__Main TS

0.3912

800.0

627.9

338.5

---

---

7__Main TS

6.266

800.0

632.1

339.0

---

---

8__Main TS

11.25

800.0

638.2

343.2

---

---

Reboiler

14.80

800.0

---

349.3

---

288.9

Column Profiles Energy Temperature (C)

Liquid Enthalpy (kJ/kgmole)

Vapour Enthalpy (kJ/kgmole)

Heat Loss (kJ/h)

Condenser

-37.85

-1.021e+005

-8.875e+004

---

1__Main TS

-32.80

-1.056e+005

-8.925e+004

---

2__Main TS

-23.89

-1.104e+005

-9.052e+004

---

3__Main TS

-15.03

-1.139e+005

-9.236e+004

---

4__Main TS

-9.595

-1.155e+005

-9.381e+004

---

5__Main TS

-5.233

-1.167e+005

-9.515e+004

---

6__Main TS

0.3912

-1.180e+005

-9.715e+004

---

7__Main TS

6.266

-1.191e+005

-9.956e+004

---

8__Main TS

11.25

-1.199e+005

-1.019e+005

---

Reboiler

14.80

-1.205e+005

-1.037e+005

---

Hyprotech Ltd. Licensed to: TEAM LND

HYSYS v3.0.1 (Build 4602)

Page 3 of 3

Case Name:

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

Unit Set:

SI

Date/Time:

Wed Jun 02 20:47:02 2004

TEAM LND Calgary, Alberta CANADA

HYSYS Column Profiles Specsheet: Column Temperature Profile

Temperature vs. Tray Number 20.0

Temperature (C)

10.0

0.000

-10.0

-20.0

-30.0

-40.0 0

1

2

3

4

5

6

7

8

9

Column Temperature Profile Column Stage

Temperature (C)

Hyprotech Ltd. Licensed to: TEAM LND

Condenser

-37.85

1__Main TS

-32.80

2__Main TS

-23.89

3__Main TS

-15.03

4__Main TS

-9.595

5__Main TS

-5.233

6__Main TS

0.3912

7__Main TS

6.266

8__Main TS

11.25

Reboiler

14.80

HYSYS v3.0.1 (Build 4602)

Page 1 of 1

Case Name:

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

Unit Set:

SI

Date/Time:

Wed Jun 02 20:47:31 2004

TEAM LND Calgary, Alberta CANADA

HYSYS Column Profiles Specsheet:

Column Properties Profile Options Selected Mass basis is selected Stage

Surface Tension

Mole Weight (Vap)

Density (Vap)

Viscosity (Vap)

Therm Cond (Vap)

Heat Cap (Vap)

Mole Weight (Lt Liq)

Condenser

8.480 dyne/cm

30.19

14.38 kg/m3

7.732e-003 cP

1.598e-002 W/m-K

1.723 kJ/kg-C

30.70

1__Main TS

8.770 dyne/cm

30.70

14.24 kg/m3

7.855e-003 cP

1.629e-002 W/m-K

1.725 kJ/kg-C

32.78

2__Main TS

9.002 dyne/cm

31.88

14.18 kg/m3

8.052e-003 cP

1.680e-002 W/m-K

1.732 kJ/kg-C

35.79

3__Main TS

8.978 dyne/cm

33.47

14.35 kg/m3

8.214e-003 cP

1.723e-002 W/m-K

1.744 kJ/kg-C

38.15

4__Main TS

8.874 dyne/cm

34.68

14.58 kg/m3

8.293e-003 cP

1.746e-002 W/m-K

1.754 kJ/kg-C

39.38

5__Main TS

8.752 dyne/cm

35.79

14.83 kg/m3

8.342e-003 cP

1.761e-002 W/m-K

1.763 kJ/kg-C

40.27

6__Main TS

8.555 dyne/cm

37.41

15.24 kg/m3

8.387e-003 cP

1.777e-002 W/m-K

1.776 kJ/kg-C

41.32

7__Main TS

8.309 dyne/cm

39.36

15.79 kg/m3

8.408e-003 cP

1.789e-002 W/m-K

1.792 kJ/kg-C

42.32

8__Main TS

8.075 dyne/cm

41.22

16.35 kg/m3

8.405e-003 cP

1.796e-002 W/m-K

1.807 kJ/kg-C

43.10

Reboiler

7.895 dyne/cm

42.66

16.81 kg/m3

8.390e-003 cP

1.799e-002 W/m-K

1.819 kJ/kg-C

43.63

Stage

Surface Tension

Density (Lt Liq)

Viscosity (Lt Liq)

Therm Cond (Lt Liq)

Heat Cap (Lt Liq)

Condenser

8.480 dyne/cm

480.6 kg/m3

9.783e-002 cP

0.1187 W/m-K

2.870 kJ/kg-C

1__Main TS

8.770 dyne/cm

492.6 kg/m3

0.1044 cP

0.1168 W/m-K

2.778 kJ/kg-C

2__Main TS

9.002 dyne/cm

504.9 kg/m3

0.1110 cP

0.1134 W/m-K

2.701 kJ/kg-C

3__Main TS

8.978 dyne/cm

510.3 kg/m3

0.1143 cP

0.1102 W/m-K

2.681 kJ/kg-C

4__Main TS

8.874 dyne/cm

511.4 kg/m3

0.1146 cP

0.1082 W/m-K

2.685 kJ/kg-C

5__Main TS

8.752 dyne/cm

511.4 kg/m3

0.1142 cP

0.1067 W/m-K

2.695 kJ/kg-C

6__Main TS

8.555 dyne/cm

510.5 kg/m3

0.1130 cP

0.1047 W/m-K

2.713 kJ/kg-C

7__Main TS

8.309 dyne/cm

508.7 kg/m3

0.1113 cP

0.1027 W/m-K

2.740 kJ/kg-C

8__Main TS

8.075 dyne/cm

506.6 kg/m3

0.1095 cP

0.1011 W/m-K

2.767 kJ/kg-C

Reboiler

7.895 dyne/cm

504.9 kg/m3

0.1081 cP

9.994e-002 W/m-K

2.788 kJ/kg-C

Hyprotech Ltd. Licensed to: TEAM LND

HYSYS v3.0.1 (Build 4602)

Page 1 of 1

Case Name:

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

Unit Set:

SI

Date/Time:

Wed Jun 02 20:47:58 2004

TEAM LND Calgary, Alberta CANADA

HYSYS Column Profiles Specsheet:

Composition vs. Tray Number 1.00 Ethane (Vap) 0.900

Propane (Vap) n-Butane (Vap)

0.800

Ethane (Light) Propane (Light)

0.700

Mole Fraction

n-Butane (Light) 0.600 0.500 0.400 0.300 0.200 1.00e-001 0.000 0

1

2

3

4

5

6

7

8

9

Column Composition Profile Options Selected Fraction is selected as the composition basis

Net is selected as flow basis

Molar basis is selected Stage

Ethane (Light Liq)

Propane (Light Liq)

n-Butane (Light Liq)

Condenser

Ethane (Vap) 0.9914

0.0086

0.0000

0.9552

0.0448

0.0000

1__Main TS

0.9552

0.0448

0.0000

0.8066

0.1934

0.0000

2__Main TS

0.8710

0.1290

0.0000

0.5925

0.4075

0.0000

3__Main TS

0.7577

0.2423

0.0000

0.4239

0.5761

0.0000

4__Main TS

0.6715

0.3285

0.0000

0.3362

0.6638

0.0000

5__Main TS

0.5924

0.4076

0.0000

0.2726

0.7274

0.0000

6__Main TS

0.4765

0.5235

0.0000

0.1978

0.8022

0.0000

7__Main TS

0.3378

0.6622

0.0000

0.1268

0.8732

0.0000

8__Main TS

0.2053

0.7947

0.0000

0.0712

0.9288

0.0000

Reboiler

0.1023

0.8977

0.0000

0.0336

0.9664

0.0000

Hyprotech Ltd. Licensed to: TEAM LND

Propane (Vap)

n-Butane (Vap)

HYSYS v3.0.1 (Build 4602)

Page 1 of 1

Unit Operations Project — Design of a rectification unit

This document was prepared as a project for the Unit Operations Course Submission Date: May 2004 Contact Information: [email protected]