Design and Operation of a Distillation Column for the Binary Mixture

Design and Operation of a

Distillation Column for the

Binary Mixture:

Propane and Hydrogen Sulfide

Project Designers:

Jonathan Sherwin

Ross Starks

CHE-305-001

W. Jeffery Horne, P.E.

Bonus Design Project

April 25, 2014

Design and Operation of a Distillation Column for the Binary Mixture: Propane and Hydrogen Sulfide

2

Table of Contents Abstract ........................................................................................................................................................ 3

Diagrams .................................................................................................................................................. 4-19

T-x-y Diagrams for H2S ........................................................................................................................... 4-5

Activity Coefficients ............................................................................................................................... 6-7

McCabe-Thiele Diagrams…………………………………………………………………………………………………………….….8-19

Process Flow Diagram for optimal Distillation Column……………………………………………………………………...20

Tables……………………………………………………………………………………………………………………………………….……..21-22

Appendix…………………………………………………………………………………………………………………………………….…..23-27

Calculations……………………………………………………………………………………………………………………………………….23

Data……………………………………………………………………………………………………………………………………..………24-26

References…………………………………………………………………………………………………………………………………………27


3

Abstract

We have been tasked with the job of determining the operating conditions and costs associated with the

design and operation of a distillation column, which is to be part of a 24/7/365 industrial operation. The

feed is a binary mixture of propane (C3H8) and hydrogen sulfide (H2S). The feed is a 50% by weight

mixture of propane and hydrogen sulfide. A feed mass flow rate of 2500 kg/hr is used. Both the distillate

and bottoms products are required to be at least 90% pure, which is attainable because the binary

mixture of propane and hydrogen sulfide is not azeotropic.

Equilibrium data was attained for a range of pressures: 0.1 atm, 1.0 atm, 5.0 atm, and 10.0 atm. T-x-y

data, activity coefficients, K-values, and relative volatility for the two compounds were used to construct

Equilibrium curves. This portion of our data and diagrams was calculated and is represented by

Equations 1-9 respectively.

McCabe-Thiele diagrams were constructed for each set of equilibrium data. For each set of pressure

data, feed conditions of a bubble-point liquid, a dew-point vapor, and a 50/50 by mass mixed

vapor/liquid were evaluated. Each feed scenario had an independent q-line and the values are

represented in Table 2. The minimum reflux was determined for each scenario. A ratio of reflux to

minimum reflux within the accepted range was chosen to be 1.3, so that R = 1.3Rmin. From the McCabe-

Thiele diagrams, the number of trays was determined for each scenario. The dew-point vapor feed at

0.1 atm, 5 atm, and 10 atm showed to be the most efficient, in terms of numbers of trays. R values and

the number of trays can both be seen in Table 1 for all conditions. R values were calculated using

Equation 12 and the number of trays was extrapolated from the McCabe-Thiele diagrams.

Total condenser and partial reboiler duties were calculated for all scenarios using Equations 14 and 15

with the results represented by Table 5. To use Equations 14 and 15 we found D and B, VB, and ΔH𝑎𝑣𝑔𝑣𝑎𝑝

by

using Equation 10, Equation 13, and the NIST Webbook respectively. Each scenario was evaluated in

terms of dollars per kilogram using equation 15 with results posted in Table 6. As seen in Table 6, the

lowest average cost is $119.98 to produce 90% pure products and this is achieved with the a bubbling

point liquid feed and an operating pressure of 10atm. Therefore, we suggest these operating conditions

to achieve profit maximization with this binary mixture: Bubbling-Point Liquid Feed, Operating Pressure

of 10atm, 10 Stages, and the feed located at Stage 5. A Process Flow Diagram for the Distillation Column

for optimal conditions can be seen in Figure 21.


4

Figure 1. T-x-y Diagram for H2S at 0.1 atm.


-98

-96

-94

-92

-90

-88

-86

-84

-82

0 0.2 0.4 0.6 0.8 1

T (d

egre

es C

)

x,y

liquid H2S

vapor H2S

-65

-60

-55

-50

-45

-40

0 0.2 0.4 0.6 0.8 1

T (d

egre

es C

)

x,y

liquid H2S

vapor H2S


5



-25

-20

-15

-10

-5

0

5

0 0.2 0.4 0.6 0.8 1

T (d

egre

es C

)

x,y

liquid H2S

vapor H2S

-5

0

5

10

15

20

25

30

0 0.2 0.4 0.6 0.8 1

T (d

egre

es C

)

x,y

liquid H2S

vapor H2S


6

Figure 5. Activity Coefficients of H2S & C3H8 vs. liquid mole fractions at 0.1 atm.


0.000

0.200

0.400

0.600

0.800

1.000

1.200

1.400

0 0.2 0.4 0.6 0.8 1

Act

ivit

y C

oef

fici

ent

Liquid Mole Fraction

H2S

C3H8

0.000

0.200

0.400

0.600

0.800

1.000

1.200

1.400

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

Act

ivit

y C

oef

fici

ent


H2S

C3H8


7



0

0.2

0.4

0.6

0.8

1

1.2

1.4

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

Act

ivit

y C

oef

fici

ent


H2S

C3H8

0.000

0.200

0.400

0.600

0.800

1.000

1.200

1.400

0 0.2 0.4 0.6 0.8 1

Act

ivit

y C

oef

fici

ent


C3H8

H2S


8

Figure 9. McCabe-Thiele Diagram for a Bubble Point Vapor Feed at 0.1 atm.

The McCabe-Thiele Diagram was created from the equilibrium curve and a 45⁰ line on a

squared chart. The dotted bottoms liquid mole fraction line, xB, was drawn at 0.1. The dotted

feed liquid mole fraction line, zF, was drawn at 0.5. The dotted distillate liquid mole fraction

line, xD, was drawn at 0.9. The q-line was drawn from zF, using equation 11 and Table 2. The

Operating Line for the Minimum Rectifying Section was drawn from the intersection of the q-

line with the equilibrium curve to the intersection of xD with the 45⁰ line. The slope of the

Minimum Rectifying Section was determined and an Rmin value was calculated. An R value was

then calculated from equation 12, and the Operating Line for the Rectifying section was then

adjusted. The Operating Line for the Stripping Section was then draw from the intersection of

the Operating Line for the Rectifying Section and the q-line to the intersection of xB with the

45⁰ line. The stage lines were then stepped off from xD to xB. The number of equilibrium stages

was then counted from the number of stage lines. This method was used in all McCabe-Thiele

Diagrams, referenced in Figures 10-20.


9

Figure 10. McCabe-Thiele Diagram for a Dew Point Vapor Feed at .1 atm.


10

Figure 11. McCabe-Thiele Diagram for a 50/50 by mass Vapor/Liquid Feed at .1 atm.


11

Figure 12. McCabe-Thiele Diagram for a Bubble Point Liquid Feed at 1 atm.


12

Figure 13. McCabe-Thiele Diagram for a Dew Point Vapor Feed at 1 atm.


13

Figure 14. McCabe-Thiele Diagram for a 50/50 by mass Vapor/Liquid Feed at 1 atm.


14

Figure 15. McCabe-Thiele Diagram for a Bubble-Point Liquid Feed at 5.0 atm.


15

Figure 16. McCabe-Thiele Diagram for a Dew-Point Vapor Feed at 5.0 atm.


16

Figure 17. McCabe-Thiele Diagram for a 50/50 by mass vapor/liquid feed at 5.0 atm.


17

Figure 18. McCabe-Thiele Diagram for a Bubble-Point Liquid Feed at 10.0 atm.


18

Figure 19. McCabe-Thiele Diagram for a Dew-Point Vapor Feed at 10.0 atm.


19

Figure 20. McCabe-Thiele Diagram for a 50/50 by mass vapor/liquid feed at 10.0 atm.


20

Reflux Drum

Reflux

Total Condenser

Boilup

Partial Reboiler999912 kW/h

Bottoms

Distillate

Feed

2500 kg/h50% wt. H2S50% wt. C3H8Pi: 10.0 atmFeed Conditions: Bubble Point Liquid

1

10

5

6

1250 kg/h xD = 0.90 H2S

1250 kg/hxB = 0.10 H2S

999831 kW/h

Figure 21. Process Flow Diagram for optimal Distillation Column for Bubble Point Liquid Feed Condition

at 10.0 atm.


21

Tables

Pressure Feed Conditions Rmin R Number of Trays

0.1 atm

Bubble-Point Liquid 0.7699 1.0008 10

Dew-Point Vapor 1.7777 2.3110 8

50/50 by mass vapor/liquid 1.6666 2.1666 10

1.0 atm




5.0 atm




10.0 atm




Table 1. Rmin values, R values, and Number of Trays for respective Feed Conditions and Pressures.

Feed Condition q slope of q-line

Bubble Point Liquid 1 vertical

Dew Point Vapor 0 horizontal

50/50 by mass vapor/liquid 0.5 -1

Table 2. q values and slope of q-line, given by equation 11, for respective Feed Conditions.

P=.1atm BP DP 50-50

m of strip 1.500 1.763 1.662

Vb 2.001 1.311 1.510

P=1atm BP DP 50-50

m of strip 1.445 1.625 1.582

Vb 2.250 1.600 1.717

P=5atm BP DP 50-50

m of strip 1.497 1.737 1.667

Vb 2.011 1.357 1.500

P=10atm BP DP 50-50

m of strip 1.497 1.737 1.667

Vb 2.011 1.357 1.500

Table 3. Representation of equation 13 for respective Pressures.

P (atm) ΔH (kJ/kg)

0.1 466.03

1 486.16

5 434.67

10 397.73

Table 4. Average Heat of Vaporization for the Binary Mixture at the respective Pressure from Nist

Webbook.


22

BP (kW/h)

DP (kW/h)

50/50 (kW/h) BP ($/kg)

DP ($/kg)

50/50 ($/kg)

P=.1atm Qc 1165534 1928826 1466037 139.86 231.46 175.92

Qr 1165534 763763 879622 139.86 91.65 105.55

139.86 161.56 140.74 AVG

P=1atm Qc 1367107 2187720 1649077 164.05 262.53 197.89

Qr 1367107 972320 1043293 164.05 116.68 125.20

164.05 189.60 161.54 AVG

P=5atm Qc 1092715 1827585 1377317 131.13 219.31 165.28

Qr 1092803 737435 815013 131.14 88.49 97.80

131.13 153.90 131.54 AVG

P=10atm Qc 999831 1672235 1260241 119.98 200.67 151.23

Qr 999912 674751 745734 119.99 80.97 89.49

119.98 140.82 120.36 AVG

Table 5. Cost analysis representative of equations 14-16.


23

Calculations

1. Antoine Equation for C3H8: 𝑙𝑜𝑔10𝑃𝑎∗ = 6.80398 −

803.80

𝑇(℃)+246.99

2. Antoine Equation for H2S: 𝑙𝑜𝑔10𝑃𝑏∗ = 6.9937 −

768.1315

𝑇(℃)+247.09

3. 𝑃𝑎 = 𝑥𝑎𝑃𝑎∗

4. 𝑃𝑏 = (1 − 𝑥𝑎)𝑃𝑏∗

5. 𝑦𝑎 =𝑃𝑎

𝑃𝑎+𝑃𝑏

6. 𝑃𝑇 = 𝑃𝑎 + 𝑃𝑏 =76 𝑡𝑜𝑟𝑟 (0.1 𝑎𝑡𝑚), 760 𝑡𝑜𝑟𝑟 (1.0 𝑎𝑡𝑚), 3800 𝑡𝑜𝑟𝑟 (5.0 𝑎𝑡𝑚), 7600 𝑡𝑜𝑟𝑟 (10.0 𝑎𝑡𝑚)

7. K-values: 𝐾𝑖 =𝑦𝑖

𝑥𝑖

8. Relative Volatility: 𝛼𝑖𝑗 =𝐾𝑖

𝐾𝑗

9. Activity Coefficients: 𝛾𝑖 = 𝐾𝑖𝑃𝑇

𝑃𝑖∗

10. Distillation Column Material Balance: 𝐹𝑧𝐹 = 𝐷𝑥𝐷 + 𝐵𝑥𝐵

11. q-line equation: 𝑦 = (𝑞

𝑞−1) 𝑥 − (

𝑧𝐹

𝑞−1)

12. Operating Line for Rectifying Section:𝑦 = (𝑅

𝑅+1) 𝑥 + (

1

𝑅+1) 𝑥𝐷; 𝑤ℎ𝑒𝑟𝑒 𝑅 = 1.3𝑅𝑚𝑖𝑛

13. Operating Line for Stripping Section:𝑦 = (𝑉𝐵+1

𝑉𝐵) 𝑥 = (

1

𝑉𝐵) 𝑥𝐵;

𝑤ℎ𝑒𝑟𝑒 𝑉𝐵 𝑖𝑠 𝑡ℎ𝑒 𝑏𝑜𝑖𝑙𝑢𝑝 𝑓𝑟𝑜𝑚 𝑡ℎ𝑒 𝑟𝑒𝑏𝑜𝑖𝑙𝑒𝑟.

14. Condenser Duty(kW/h): 𝑄𝐶 = 𝐷 ∗ (𝑅 + 1) ∗ Δ𝐻𝑎𝑣𝑔𝑣𝑎𝑝

15. Reboiler Duty(kW/h): 𝑄𝑅 = 𝐵 ∗ 𝑉𝐵 ∗ Δ𝐻𝑎𝑣𝑔𝑣𝑎𝑝

16. $

𝑘𝑔= (

𝑘𝑊

ℎ)𝑄 ∗

.09$

(𝑘𝑊

ℎ)

∗1

𝐵𝑜𝑟 𝐷∗

1

60%


24

Raw Data

H2S C3H8 P (atm) torr

A 6.99387 6.80398 0.1 76

B 768.132 803.81 1 760

C 247.09 246.99 5 3800

10 7600

xp xh T (C) Pvapp (Torr) Pvaph (Torr) Pp (Torr) Ph (Torr) P (Torr) yp yh Kp Kh A12 A21 gam12 gam21

0 1 -96.861 28.175 76.000 0.000 76.000 76.000 0.000 1.000 0.371 1.000 0.371 2.697 0.238 1.000

0.05 0.95 -96.452 29.134 78.467 1.457 74.544 76.000 0.019 0.981 0.383 1.032 0.371 2.693 0.308 0.997

0.1 0.9 -96.028 30.157 81.094 3.016 72.985 76.000 0.040 0.960 0.397 1.067 0.372 2.689 0.381 0.988

0.15 0.85 -95.588 31.250 83.897 4.688 71.313 76.000 0.062 0.938 0.411 1.104 0.372 2.685 0.455 0.977

0.2 0.8 -95.132 32.421 86.895 6.484 69.516 76.000 0.085 0.915 0.427 1.143 0.373 2.680 0.529 0.963

0.25 0.75 -94.656 33.678 90.107 8.420 67.581 76.000 0.111 0.889 0.443 1.186 0.374 2.676 0.602 0.947

0.3 0.7 -94.161 35.031 93.558 10.509 65.491 76.000 0.138 0.862 0.461 1.231 0.374 2.671 0.672 0.930

0.35 0.65 -93.644 36.490 97.275 12.772 63.228 76.000 0.168 0.832 0.480 1.280 0.375 2.666 0.740 0.913

0.4 0.6 -93.104 38.069 101.287 15.228 60.772 76.000 0.200 0.800 0.501 1.333 0.376 2.661 0.805 0.896

0.45 0.55 -92.539 39.782 105.633 17.902 58.098 76.000 0.236 0.764 0.523 1.390 0.377 2.655 0.866 0.878

0.5 0.5 -91.946 41.647 110.353 20.823 55.177 76.000 0.274 0.726 0.548 1.452 0.377 2.650 0.924 0.861

0.55 0.45 -91.323 43.684 115.498 24.026 51.974 76.000 0.316 0.684 0.575 1.520 0.378 2.644 0.978 0.844

0.6 0.4 -90.668 45.916 121.126 27.550 48.450 76.000 0.362 0.638 0.604 1.594 0.379 2.638 1.029 0.828

0.65 0.35 -89.977 48.374 127.306 31.443 44.557 76.000 0.414 0.586 0.636 1.675 0.380 2.632 1.076 0.812

0.7 0.3 -89.245 51.091 134.121 35.764 40.236 76.000 0.471 0.529 0.672 1.765 0.381 2.625 1.119 0.797

0.75 0.25 -88.470 54.109 141.673 40.582 35.418 76.000 0.534 0.466 0.712 1.864 0.382 2.618 1.160 0.782

0.8 0.2 -87.645 57.479 150.083 45.983 30.017 76.000 0.605 0.395 0.756 1.975 0.383 2.611 1.198 0.767

0.85 0.15 -86.766 61.264 159.502 52.075 23.925 76.000 0.685 0.315 0.806 2.099 0.384 2.604 1.233 0.753

0.9 0.1 -85.824 65.543 170.117 58.988 17.012 76.000 0.776 0.224 0.862 2.238 0.385 2.596 1.265 0.740

0.95 0.05 -84.811 70.413 182.163 66.893 9.108 76.001 0.880 0.120 0.926 2.397 0.387 2.587 1.295 0.728

1 0 -83.719 76.000 195.933 76.000 0.000 76.000 1.000 0.000 1.000 2.578 0.388 2.578 1.300 0.715


0 1 -60.336 314.465 760.000 0.000 760.000 760.000 0.000 1.000 0.414 1.000 0.414 2.417 0.289 1.000

0.05 0.95 -59.748 324.409 782.926 16.220 743.780 760.000 0.021 0.979 0.427 1.030 0.414 2.413 0.355 0.998

0.1 0.9 -59.139 334.964 807.226 33.496 726.504 760.000 0.044 0.956 0.441 1.062 0.415 2.410 0.423 0.992

0.15 0.85 -58.509 346.188 833.027 51.928 708.073 760.001 0.068 0.932 0.456 1.096 0.416 2.406 0.492 0.983

0.2 0.8 -57.855 358.143 860.465 71.628 688.372 760.001 0.094 0.906 0.471 1.132 0.416 2.403 0.561 0.972

0.25 0.75 -57.176 370.900 889.701 92.725 667.275 760.000 0.122 0.878 0.488 1.171 0.417 2.399 0.629 0.960

0.3 0.7 -56.470 384.541 920.911 115.362 644.638 760.000 0.152 0.848 0.506 1.212 0.418 2.395 0.695 0.947

0.35 0.65 -55.736 399.158 954.300 139.705 620.295 760.000 0.184 0.816 0.525 1.256 0.418 2.391 0.759 0.933

0.4 0.6 -54.970 414.856 990.096 165.943 594.058 760.000 0.218 0.782 0.546 1.303 0.419 2.387 0.821 0.919

0.45 0.55 -54.172 431.756 1028.563 194.290 565.710 760.000 0.256 0.744 0.568 1.353 0.420 2.382 0.880 0.905

0.5 0.5 -53.337 449.996 1070.005 224.998 535.002 760.000 0.296 0.704 0.592 1.408 0.421 2.378 0.936 0.890

0.55 0.45 -52.463 469.735 1114.768 258.354 501.646 760.000 0.340 0.660 0.618 1.467 0.421 2.373 0.990 0.876


25

0.6 0.4 -51.547 491.161 1163.259 294.697 465.304 760.000 0.388 0.612 0.646 1.531 0.422 2.368 1.040 0.862

0.65 0.35 -50.585 514.490 1215.948 334.418 425.582 760.000 0.440 0.560 0.677 1.600 0.423 2.363 1.088 0.848

0.7 0.3 -49.572 539.976 1273.388 377.984 382.017 760.000 0.497 0.503 0.710 1.676 0.424 2.358 1.133 0.835

0.75 0.25 -48.503 567.923 1336.231 425.942 334.058 760.000 0.560 0.440 0.747 1.758 0.425 2.353 1.175 0.821

0.8 0.2 -47.374 598.688 1405.250 478.950 281.050 760.000 0.630 0.370 0.788 1.849 0.426 2.347 1.215 0.808

0.85 0.15 -46.177 632.700 1481.368 537.795 222.205 760.000 0.708 0.292 0.833 1.949 0.427 2.341 1.252 0.796

0.9 0.1 -44.906 670.479 1565.700 603.431 156.570 760.001 0.794 0.206 0.882 2.060 0.428 2.335 1.287 0.784

0.95 0.05 -43.551 712.654 1659.591 677.021 82.980 760.001 0.891 0.109 0.938 2.184 0.429 2.329 1.320 0.772

1 0 -42.102 760.000 1764.697 760.000 0.000 760.000 1.000 0.000 1.000 2.322 0.431 2.322 1.360 0.761


0 1 -22.101 1697.169 3800.000 0.000 3800.000 3800.000 0.000 1.000 0.447 1.000 0.447 2.239 0.325 1.000

0.05 0.95 -21.296 1747.730 3908.015 87.386 3712.614 3800.001 0.023 0.977 0.460 1.028 0.447 2.236 0.388 0.998

0.1 0.9 -20.464 1801.199 4022.089 180.120 3619.880 3800.000 0.047 0.953 0.474 1.058 0.448 2.233 0.453 0.994

0.15 0.85 -19.602 1857.828 4142.736 278.674 3521.326 3800.000 0.073 0.927 0.489 1.090 0.448 2.230 0.518 0.987

0.2 0.8 -18.710 1917.894 4270.527 383.579 3416.421 3800.000 0.101 0.899 0.505 1.124 0.449 2.227 0.583 0.978

0.25 0.75 -17.785 1981.709 4406.097 495.427 3304.573 3800.000 0.130 0.870 0.522 1.159 0.450 2.223 0.648 0.968

0.3 0.7 -16.824 2049.625 4550.161 614.888 3185.113 3800.000 0.162 0.838 0.539 1.197 0.450 2.220 0.711 0.957

0.35 0.65 -15.826 2122.034 4703.520 742.712 3057.288 3800.000 0.195 0.805 0.558 1.238 0.451 2.217 0.772 0.945

0.4 0.6 -14.788 2199.382 4867.079 879.753 2920.247 3800.000 0.232 0.768 0.579 1.281 0.452 2.213 0.832 0.933

0.45 0.55 -13.706 2282.173 5041.859 1026.978 2773.022 3800.000 0.270 0.730 0.601 1.327 0.453 2.209 0.889 0.921

0.5 0.5 -12.578 2370.979 5229.021 1185.489 2614.511 3800.000 0.312 0.688 0.624 1.376 0.453 2.205 0.945 0.909

0.55 0.45 -11.400 2466.455 5429.889 1356.550 2443.450 3800.000 0.357 0.643 0.649 1.429 0.454 2.201 0.998 0.896

0.6 0.4 -10.168 2569.350 5645.974 1541.610 2258.390 3800.000 0.406 0.594 0.676 1.486 0.455 2.197 1.048 0.884

0.65 0.35 -8.878 2680.529 5879.017 1742.344 2057.656 3800.000 0.459 0.541 0.705 1.547 0.456 2.193 1.096 0.872

0.7 0.3 -7.523 2800.990 6131.024 1960.693 1839.307 3800.000 0.516 0.484 0.737 1.613 0.457 2.189 1.142 0.860

0.75 0.25 -6.100 2931.892 6404.323 2198.919 1601.081 3800.000 0.579 0.421 0.772 1.685 0.458 2.184 1.186 0.848

0.8 0.2 -4.601 3074.593 6701.629 2459.674 1340.326 3800.000 0.647 0.353 0.809 1.764 0.459 2.180 1.227 0.836

0.85 0.15 -3.018 3230.684 7026.124 2746.081 1053.919 3800.000 0.723 0.277 0.850 1.849 0.460 2.175 1.266 0.825

0.9 0.1 -1.345 3402.049 7381.558 3061.844 738.156 3800.000 0.806 0.194 0.895 1.943 0.461 2.170 1.304 0.814

0.95 0.05 0.430 3590.927 7772.385 3411.381 388.619 3800.000 0.898 0.102 0.945 2.045 0.462 2.164 1.339 0.803

1 0 2.316 3800.000 8203.922 3800.000 0.000 3800.000 1.000 0.000 1.000 2.159 0.463 2.159 1.380 0.792


0 1 -0.345 3507.561 7600.000 0.000 7600.000 7600.000 0.000 1.000 0.462 1.000 0.462 2.167 0.340 1.000

0.05 0.95 0.597 3609.151 7810.044 180.458 7419.542 7599.999 0.024 0.976 0.475 1.028 0.462 2.164 0.402 0.998

0.1 0.9 1.571 3716.406 8031.510 371.641 7228.359 7600.000 0.049 0.951 0.489 1.057 0.463 2.161 0.465 0.994

0.15 0.85 2.578 3829.794 8265.330 574.469 7025.531 7600.000 0.076 0.924 0.504 1.088 0.463 2.158 0.529 0.988

0.2 0.8 3.621 3949.841 8512.540 789.968 6810.032 7600.000 0.104 0.896 0.520 1.120 0.464 2.155 0.592 0.980

0.25 0.75 4.702 4077.128 8774.291 1019.282 6580.718 7600.000 0.134 0.866 0.536 1.155 0.465 2.152 0.655 0.971

0.3 0.7 5.823 4212.307 9051.867 1263.692 6336.307 7599.999 0.166 0.834 0.554 1.191 0.465 2.149 0.717 0.961

0.35 0.65 6.988 4356.109 9346.711 1524.638 6075.362 7600.000 0.201 0.799 0.573 1.230 0.466 2.146 0.778 0.950

0.4 0.6 8.199 4509.352 9660.431 1803.741 5796.259 7600.000 0.237 0.763 0.593 1.271 0.467 2.142 0.837 0.939

0.45 0.55 9.459 4672.964 9994.848 2102.834 5497.166 7600.000 0.277 0.723 0.615 1.315 0.468 2.139 0.893 0.928

0.5 0.5 10.772 4847.989 10352.010 2423.995 5176.005 7600.000 0.319 0.681 0.638 1.362 0.468 2.135 0.948 0.916

0.55 0.45 12.142 5035.617 10734.250 2769.589 4830.411 7600.000 0.364 0.636 0.663 1.412 0.469 2.132 1.001 0.905

0.6 0.4 13.574 5237.200 11144.200 3142.320 4457.680 7600.000 0.413 0.587 0.689 1.466 0.470 2.128 1.052 0.893

0.65 0.35 15.073 5454.289 11584.890 3545.288 4054.712 7600.000 0.466 0.534 0.718 1.524 0.471 2.124 1.100 0.881


26

0.7 0.3 16.643 5688.662 12059.790 3982.063 3617.937 7600.000 0.524 0.476 0.749 1.587 0.472 2.120 1.146 0.870

0.75 0.25 18.292 5942.374 12572.880 4456.780 3143.220 7600.000 0.586 0.414 0.782 1.654 0.473 2.116 1.190 0.859

0.8 0.2 20.026 6217.805 13128.780 4974.244 2625.756 7600.000 0.655 0.345 0.818 1.727 0.474 2.111 1.233 0.847

0.85 0.15 21.853 6517.731 13732.860 5540.071 2059.929 7600.000 0.729 0.271 0.858 1.807 0.475 2.107 1.273 0.837

0.9 0.1 23.782 6845.402 14391.390 6160.861 1439.139 7600.001 0.811 0.189 0.901 1.894 0.476 2.102 1.311 0.826

0.95 0.05 25.824 7204.645 15111.750 6844.412 755.588 7600.000 0.901 0.099 0.948 1.988 0.477 2.098 1.347 0.816

1 0 27.989 7600.000 15902.660 7600.000 0.000 7600.000 1.000 0.000 1.000 2.092 0.478 2.092 1.370 0.806


27

References

Seader, J. D., Ernest J. Henley, and D. Keith Roper. Separation Process Principles. Third ed. N.p.:Courier Westford, 2011. Print.

NIST Chemistry Webbook. N.p., n.d. Web. 24 Apr. 2014. <http://webbook.nist.gov/chemistry/>. Thermophysical Properties of Fluid Systemsby E.W. Lemmon, M.O. McLinden, D.G. Friend

http://www.noodletools.com/noodlebib/citeone_s.php?id=4972639

http://www.noodletools.com/noodlebib/citeone_s.php?id=4972639

Documents

Design and Operation of a Distillation Column for the Binary Mixture