Data Book on Hydrocarbons
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DATA BOOK ON
HYDROCARBONS APPLICA TION TO PROCESS ENGINEERING
by
J. B. MAXWELL NINTH PRINTING
ROBERT E. KRIEGER PUBLISHING COMPANY MALABAR, FLORIDA
•
ORIGINAL EDITION 1960 REPRINTED 1977 FROM NINTH PRINTING 1968
Printed and Published by ROBERT E. KRIEGER PUBLISHING COMPANY, INC. KRIEGER DRIVE MALABAR, FLORIDA 32950
© Copyright 1950 by
STANDARD OIL DEVELOPMENT COMPANY Reprinted by Arrangement with VAN NOSTRAND REINHOLD CoMPANY
All rights reserved. No reproduction in any form of this book, in while or in part (except for brief quotation in critical articles or reviews), may be made without written authorization frOm the publisher.
PRINTED IN THE UNITED STATES OF AMERICA
Library of Congress Cataloging in Publication Data
Maxwell, J B 1902Data book on hydrocarbons. Reprint of the 9th printing published in 1968 by Van Nostrand, Princeton, N. J., in The Esso series. Includes bibliographies. 1. Hydrocarbons. I. Title. TP690.M35 1975 661'.81 74-30163 ISBN 0-88275-257-X
PREFACE The primary purpose of this book is to provide (1) basic data on hydrocarbons and petroleum fractions, (2) methods of applying these data to process engineering, including illustrative examples and some fundamental theory, and (3) applications of a few of the unit operations of chemical engineering uscd extensively in the petroleum industry. Earlier editions of the present volume have been used in the Standard Oil Development Company and other affiliates of the Standard Oil Company (New .Jersey). Because this book has proved to be quite valuable to technical personnel, the Standard Oil Development Company has decided to make it available for practicing engineers and students of petroleum technology. The author is very much indebted to many associates in the preparation of thi s book and, in particular, to W. H. Hatch for invaluable assistance in editing the text and preparing the charts for publication, to C. O. Rbys, Sr., for the derivation of the .Mollier diagrams and other charts, to C. J. Robrecht (or constructive criticism and advice during the preparation of the manuscript. Furthermore, any list of acknowledgments would be incomplete without mentioning R. S. Piroomov who was responsible for the early development of a company data book. J. B. MAXWELL Standard Oil Development Company Linden, New Jersey
•
CONTENTS PHYSICAL DATA SElCTIOl\
PAGE
1. PHYSICAL CONSTANTS.......................................
1
Hydrocarbons, 2-Miscellaneous Organic Compounds, 6--MisceIlaneous Gases, 9 2. CHARACTERISTICS OF PETROLEUM FRACTIONS... . . . . . . . ..
10
Average Boiling Point, 14-Characterization Factor, Hi-Gravity, 18 3. MOLECULAR WEIGHT. . . . . . . . . . . . . . . . . .. . . . . .. ... ... . . . . . . . . .
19
Paraffins, 20-Petroleum Fractions, 21 4. VAPOH PRESSURE
.
24
Paraffins and Olefins, 27-Diolefins and Acetylenes, 35-Aromatics, 37 -Cycloparaffins, 39-Hydrocarbons, 40-Gasolines, 44 5. FUGACITY
.
45
Fugacity Function of Individual Hydrocarbons, 49-Fugacity Function of Hydrogen, Ol-Fugacity of Hydrocarbon Vapors, 62-Relative Volatility of LigM Hydrocarbons, 6~-Fugacity Correction Factor for Light Hydrocarbons in Absorber Oib, 67 6. CRITICAL PROPERTIES.......................................
68
Critical Temperature of Pure Hydrocarbons, 69-Critical Temperature of Light Hydrocarbon Mixtures, 'i'O-Critical Pressure of K ormal Paraffins, 71-Critical Temperature and Pressure of Petroleum Fractions, 72 7. THEHMAL PHOPERTlES
.
Specific Heats of Gases and Vapors, 88-Enlhalpy-Presoure Relationship for Hydrocarbon Vapors, 92-Bpecifjr Heats of Liquid Hydrocarbons and Petroleum Fractions, 93-Latenl Ileat of Vaporization of Light Hydrocarbons and Normal Paraffins, 94-Enthalpy of Individual Hydrocarbons, 98-Enthalpy of Petroleum Fractions, 114-Mollier Diagrams for Light Hydrocarbons, 128 VII
75
viii
CONTENTS PAGE
SECTION
8. DENSITy........................... . . . . . . . . . . . . . . . . . . . . . . . . . . .
136
Conversion Charts for 0 API Gravity, 138-Specific Gravity of Saturated Hydrocarbon Liquids, 14o-Thermal Expamlion of Liquid Petroleum Fractions, 143-P-V-T Relations of Hydrocarbon Vapors, 148 9. VISCOSITY. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .
155
Conversion Charts, 158-Viscosity of Hydrocarbons and Crude Fractions, 161-Viscosity-Temperature Charts, 166--Viscosity Index of Lubricating Oils, 168-Viscosity Blending Index, 173-Viscosity of Hydrocarbon Vapors and Miscellaneous Gases, 174 10. COMBUSTION. . .
178
Heat of Combustion of Petroleum Fractions and Hydrocarbon Gases, 18o-Enthalpy of Flue Gas Components, 182-Heat Available from the Combustion of Refinery Gases and Fuel Oils, 184-Properties of Flue Gases, 189 UNIT OPERATIONS 11. FLOW OF FLUIDS.............................................
193
Friction Factor for Fluid Flow, 19B-Pressure Drop in Commercial Pipes, 199-Equivalent Length of Fittings, 202-Friction Loss Due to Contraction and Enlargement, 204-Discharge Characteristics of Weirs, 205-Pressure Drop Across Tube Banks, 206 12. FLOW OF HEAT. . . . . . . . . . . . . . . . . . . . . . . . . . .. . .
.. .
207
Heat Loss by Radiation and Natural Convection, 209-Heat Transfer to Fluids Inside Tubes, 211-Heat Transfcr to Fluids Outside Tubes, 212-Thermal Conductivity of Petroleum Fractions, Water, and Gases, 213-Logarithmic Mean Temperatme Difference, 217 13. EQUILIBRIUM FLASH VAPORIZATION.. . . . . . . . . . . . . . . . . . . . . ..
222
14. FRACTIONATING TOWERS.. ..
230
Minimum Reflux Ratio and Theoretical Steps, 23O-Correlation of Theoretical Steps with Reflux Ratio, 244-0verall Plate Efficiency, 245-Packed Towcrs, 246 CONVERSIOX FACTORS. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . ..
249
INDEX............................................................
253
Section I
PHYSICAL CONSTANTS In the following tables the more common physical constants are given for hydrocarbons, certain other organic series, and miscellaneous gases. While these constants, in general, are based upon reliable data, estimated "alues were included in a few instances where available data were considered questionable. Where no reasonably good basis was available for either estimating or calculating the constants, they are omitted. The density, boiling point, melting point, and heat of eo'mbustion for most of the hydrocarbons are those given in the Burea1t of Standards Circular C461. 1 GENERAL REFERENCES Annual Tables of Physical Constants, Nat. Research Council (19-11, 1942). Beattie, Poffenberger and Hadlock, J. Chem. Phys. 3, 96 (1935). Beattie, Simard and Su, J. Alii. Chem. Soc. 61, 24 (1939); 61,924 (1939). Cole and Cole, J. Chem. Phys. 9, 341 (1941). Doss, "Physical ~nstants of the Principal Hydrocarbons," 4th Edition, The Texas Co., New York, N.Y. (1943). Ginnings, J. Am. Cltell/. Soc. 62, 1923 (1940). Ginnings and Baum, J. Am. Chem. Soc. 59, 1111 (1937). Ingersoll, Thesis, ~Iass. Inst. Tech. (1930). International Critical Tables, Vols. I and III. Kay, Ind. Eng. Chem. 30, 459 (1938). Kharasch, J. Research Nat. Bur. Standards 2,359 (1929). Krase and Goodman, Ind. Eng. Chelll. 22, 13 (1930). Meyers, Scott, Brickwede and RAnds, Unpublished Data, Nat. Bur. Standards. Pickering, Bur. Standards Sci. Paper 511 (1926). Rintelen, Gross and Saylor, J. Am. Chelll. Soc. 62, 1923 (19-10). Tables anntlelles de wnstantes et dunnee nUllteriqlte, Vols. VII to XIII (1925-1939). I'
"Sclcdcd Values of Propertips of Hydrocarbons," Nal. Bur.
(947).
1
~lalldards
Circular Cl,61
~
PHYSICAL CONSTANTS OF HYDROCARBONS
. FOR~tULA
MOLEC. WT.
BOILING POINT of
:.fELTING POINT of
NORMAL PARAFFINS Methane ................... Ethane .................... Propane ................... Bu~no ....................
CH, C,H. C,H, C,H,o
Pentane ................... Hexane .................... Heptane ................... Octane ...... '" ...........
C,H 12 C,H" C,H 16 C,H,s
72.1 86.2 100.2 114.2
96.9 -201. 5 155.7 -139.5 209.2 -131. 1 258.2 - 70.3
Nonane .................... Decane .................... Undecane .................. Dodecane .... , .............
C,H,o C1oH" C"H,. C12 H "
128.2 142.3 156.3 170.3
303.4 345.2 384 .4 421.3
ISO-PARAFFINS Isobutane ..................
C,H 1o
58.1
10.9 -255.0
C,H 12
72.1
82.2 -255.5
C,H"
72.1
49.0
2-Methylpentanc (Isohexane) , 3-Methylpcnt,ane, ........ 2,2-Dimethylbutane ();eohexane) . ' .. , ......... , ... 2,3-Dimethylbutane (Diisopropyl) ...............
C,H" C,H"
86.2 86.2
140.5 -245 145.9 -180
C,H"
86.2
C.H"
86.2
2-Met,hylhexane (Isoheptane) . 3-Methylhexanc ............ 3-Ethylpentane .... ...... 2,2-Dimethylpentane ....... '
C,H I6 C,H" (',H" (',H"
2,3·Dimethylpentane ........ 2,4-Dimethylpentane ........ 3,3-Dimethylpentane ....... ,
C,H" C,H" C,H"
2-Methylbutane (Isopentane), 2,2-Dimeth:dpropane (~eopentane). . ..............
16.0 -258.9 -296.5 30.1 -128.0 -297.8 44.1 - 43.8 -305.7 58.1 31.1 -216.9
DENSITY
°API
Sp Gr 60°/60° Lb/gal
t
P Atm
G/ml
206.3 306
45.8 48.2 42.0 37.4 32.6 29.4 26.8 24.6
of
D
HEAT OF COMBUSTION
@
60°F-BTU /Ib
Gross
Net
0.162 .203 .226 .225
23,860" 22,300" 21,650" 21,290'
21,500" 20,42019,930" 19,670"
.232 .234 .234 .233
21,070" 20,780 20,670 20,590
19,500" 19,240 19,160 19,100
20,530 20,480 20,450 20,420
19,050 19,020 19,000 18,980
0.30 .374 .508 .584
2.50 3.11 4.23 4.86
92.7 81.6 74.2 6ti,{j
.631 .664 .688 .707
5.25 5.53 5.73 5.89
386.5 455.0 512.5 565
64.5 61.3 58.7 56.4
. -.).) ,- ... .734 .744 , 753
6.01 6.11 6.19 6.27
612' 654' 695' 731"
.563
4.69
275
36
.234
21,240'
19,610"
.625
5.20
369.5
32.4
.234
21,030"
19,450"
.597
4.97
329"
35'
-
20,960'
19,330"
o;.:-
83.5 80.0
.658 , 669
5.48 5.57
437' 443'
31' 30"
-
-
20,750 20,760
19,210 19,220
tp
121.5 -147 6
84.9
.654
5.44
415'
31'
-
20,700
19,160
136.4 -198.8
81.0
.666
5.54
441
31
20,740
19,200
100.2 100.2 100.2 100.2
194.1 197.5 200.2 174.6
-180.8 -182.9 -181.5 -190.8
75.i 73.0 69.ti 77.'2
. 68:l .692 .703 .678
5.68 5.76 5.85 5.64
496 504 508' 475'
28" 28.5" 28.5 28.5
20,650 20,660 20.670 20,600
19,140 19,150 19,160 19,090
100.2 100.2 100.2
193.6 176.9 -183.1 186.9 -211.0
-') , I ._
70.6
.700 .678 .698
5.83 5.54 5.81
498' 477 487'
29 28.5' 28'
20,540 20,620 20,620
19,130 19,110 19,110
+
+
64.5 21.5 14.1 14.7
340 247 147 111
CRITICAL CONSTANTS
-116.3
+ 90.1
23" 22"
20" lb'
-
-
d
;.:>-3
;.:tp
o o
~
o Z
:I:
+
2.1
120 94.9 105
71.2
..
. 241
-
-
-
-
-
>
Z
~ lfJ
-
-
20,150 Q
-
-
20,320Q -
19,210Q
-
-
20,060"
18,950Q
...
;> t"'i
19,180Q
19,040" 19,04QQ
* Mixture of cis- and ** Sublimes.
-
~ c; o o ~
o
Z ::t: ~
-
-114
9.8 -153
+
47.7 -188.5 80.4 - 26.0
104.4 -159 132.8 -148
209
60~
Pentyne-1 (Propylacetylene). Pentyne-2 .................. 3-Methylbutyne-1 (Isopropylacetylene) ................
CsH s CsH s
68.1 68.1
CsHs
68.1
79.7
.670
5.58
410"
Hexyne-1 (Butylacetylene) ... Hexyne-2 .................. Hexyne-3 ...............•..
CsH lo CSH lO CsH lO
82.1 82.1 82.1
160.9 -205.6 184.1 -126.4 179.2 -149.8
65.0 60.8 63.1
.720 .736 .727
5.99 6.13 6.05
-
-
4-Methylpentyne-1 .......... 4-Methylpentyne-2 .......... 3,3-Diroethylbutyne-1 .......
CSH lO CSH lO CaRlO
82.1 82.1 82.1
142.1 -157.1 162 100.0 -114.2
67.5 65.3 78.7
.711 .719 .673
5.92 5.98 5.60
-
-
C.H.
52.1
73.9
.689
5.73
365"
OLEFINS-ACETYLENES Buten-3-yne-1 (VinylAcetylene) .................•..
D G/rol
HEAT OF COMBUSTION @ 60°F-BTU lib
82
42
-
-
20,550
20,74~
19,46()Q
-
-
-
-
-
75.;
-
-
-
-
-
-
tj
~
o o > ~ c;
C
Z
en
-
58.7
0.744
6.19
49.4
.782
6.51
-
56.4 32.8
.753 .861
6,27 7.17
-
-
176.2
41.9
28.6
.884
7.36
231.1
-139.0
30.8
.872
7.26
106.2 106.2 106.2 106.2
292.0 - 13.3 282,4 - 54.2 281.0 55.9 277 .1 -138.9
28.4 31.3 31.9 30.8
.885 .869 .866 .872
C,Hu
120.2
349.0 -
13.8
25.7
C,Hu
120.2
336.5 -
47.3
C,H"
120.2
C9H l2 C9Hn C,H" C,H" C,R"
120.2 120.2 120.2 120.2 120.2
318.6 306.3 329.2 322.7 324.5
CYCLOPARAFFINS Cyclopropane ... ............
C,H.
42.1
-
Cyclobutane . ...............
C.H.
56.1
+ 54.7
Cyclopentane . ...•.......... Methylcyc1opentane . ........ 1, I-Dimethylcyclopentane ....
CeHlo CtHn C,H u
l,2··)imethylcyc1opentane-cis. 1,2-Dimethylcyclopentanetrans . ................... 1,3-Dimethylcyclopentanetrans . ................... Ethylcyclopentane ..........
Cyclohexane ................ Methylcyclohexane . .........
Penten-l-yne-3 . ............ (AllylacetyPenten-l-yne-4 lene) . ................... 2-Methylbuten·l-yne·3 . .....
C.He
66.1
138.6
C,H e CoHo
66.1 66.1
107 90
Hexen-l-yne-3 . ............. Hexen-l-yne-5 .. ............ 2-Methylpenten-l-yne-3 ..... 3--Methylpenten-3-yne-l* .....
C.H. CoHo CoHo CoHo
80.1 80.1 80.1 80.1
185 158 169 156
AROMATICS Benzene ... .................
CoHo
78,1
Toluene ..... ...............
C,H.
92.1
o-Xylene .... ............... m-Xylene .........•....... . ;p-Xylene . .............. , ... Ethylbenzene .... ...........
C.H IO CaHlo CaHlo CaH lo
1,2,3-Trimethylbenzene . ..... 1,2,4-Trimethylbenzene (Pseu· documene) . .............. 1,3,5·Trimethyl benzene (Mesitylene) . ................ Propylbenzene . ............. lsopropylbenzene (Cumene) .. I-Methyl-2-Ethylbenzene .... I-Methyl-3-Ethylbenzene .... I-Methyl-4-Ethylbenzene ....
-
-
-
-
-
-
-
-
-
-
-
-
---
-
--
-
551.3
47.9
0.304
17,990
17.270
609.1
41.6
.292
18,270
17,450
7,37 7.24 7.21 7.26
675 655· 652 655
37 36' 35' 38
.288· .288· .270'
-
18.500 18,500 18,430 18,490
17,610 17,610 17,540 17.600
.900
7.49
72C/'
32'
-
-
-
29.1
.881
7.34
708'
33
-
18,570
17.620
328.3 - 48.6
31.1
.870
7.24
700'
-
18,620
17,670
-147,1 -140.8 -126,6
.866 .866 .883 .870 .868
7,21 7.21 7.35 7.24 7.23
690 68C/' 702< 695' 696'
-
18,660 18.670
17,710 17,720
o
34' 34'
82.7
31.9 31.9 28.7 31.1 31.5
33 /' 34' 35'
27.0 -196.6
98.6
.615
5.12
256
54
-
-
58.0
74.8
.686
5.71
385'
50'
-
cB
70.1 84.2 98.2
120,7 -136.7 161.3 -224.4 189.5 -105
56,9 56.2 54.7
.751 .754 .760
6.25 6.28 6.33
470' 520' 550'
46'
C7H u
98.2
210.7 -
62
50.4
.778
6.48
570'
40'
C,H u
98.2
197.4 -182
65.4
.757
6.30
560'
41'
-
C 7H 14 C 7H u
98.2 98.2
195.4 -213 218.2 -217
57.2 52.0
.750 .771
6.24 6.42
555' 580'
41' 40'
-
20,110
18.760
CeRn C 7H 14
84.2 98.2
177.3 44 213.6 -195.6
49.0 61.3
.784 .774
6,53 6.44
538 575
40.4 40'
.273
20.030 20.000
18.680 18.650
+
, Heat of combustion &8 a gas-otherwise &8 a liquid. • Estimated.
-
-
-
+
-
-
-
-
Critical temperature-boiling point correlatioil. ., Vapor pressure curve or correlation.
C
34'
42' 42'
-
-
-
-
-
-
--
-
-
-
-
-
20,350' 20,110
18.760
20,020
18,670
20.020
18,670
-
-
"'d
~ ~
......
o
~
19.ססOO
-
-
• Mixture of cis- and trans' isomers. •• Sublimes.
o:n
PHYSICAL CONSTANTS OF ORGANIC COMPOUNDS 0:.
FORMULA
MOLEC. WT.
BOILING POINT of
MELTING POINT of
CRITICAL CONSTANTS
DENSITY
Sp Gr
600 j600 Lb jgal
t
of
HEAT OF COMBUSTION HEAT OF VAPORIZ.
@B.P. P D Atm Gjml BTU lib
@ 60°F-BTU /lb Gross
Net
-- -ALCOHOLS Methanol (Methyl Alcohol) .. CHaOH Ethanol (Ethyl Alcohol) ..... CH aCH20H Propanol-1 (Normal Propyl Alcohol) ................. CH aCH2CH20H Propanol-2 (Isopropyl Alcohoi) ........ , ............ (CHa),CHOH Butanol-1 (Normal Butyl Alcohol) ................. Butanol-2 (Sec. Butyl Alcohol) 2-Methylpropanol-l (Isobutyl Alcohol) ................. 2-Methylpropanol-2 (Tert. Butyl Alcohol) ........... Pentanol-1 (Normal Amyl Alcohol) ................. Pentanol-2 (Sec. Amyl Alcohoi) ..................... Pentanol-3 (Diethyl Carbinol) 2-Methylbutanol-l (Sec. Butyl Carbinol) ................ 2-Methylbutanol-2 (Tert. Amyl Alcohol) ........... 3-Methylbutanol-1 (Isoamyl Alcohol) ........ , ........ 3-Methylbutanol-2 (Methyl Isopropyl Carbinol) ....... 2.2-Dimethylpropanol-l (Tert. Butyl Carbinol) ..........
474
9760
8580
.794
6.61 469.6 63.1
.275
361
12,780
11,550
207.0 -195
.808
6.73 506.7 50.0 .273
296
14,450
13,190
180.2 -129
.789
6.57
289
14.350
13.090
> ~
148.1 -143.7 0.796
46.1
173.0 -174
60.1 60.1
tj
254 242
15.500
14,220
-
-
-
-
-
249
15.450
14.170
o
(6.60)
-
-
-
235
15,290
14,010
.819
6.82
-
-
223*
16,220
14,930
.814 .826
6.78 6.88
-
-
-
213* 211*
-
...
.820 6.83 - .825 -6.87
-
-
-
218*
-
-
77.9
(.793)
280.4 -109.8
(CH ahCHCH 2OH
74.1
226.4 -162
(CHa)aCOH
74.1
180.7
CH a(CH 2) aCH 20H
88.1
-
-
CH a(CH2hCH(OH)CH a (CH.CH 2hCHOH
88.1 88.1
247.1 240
-
CH aCH2CH(CH a)CHIOH
88.1
264
-
to
-
6.71
-129.6
-
-
.806
243.9 211.1
-
o o
6.78 549 6.75
74.1 74.1 ,
-
.814 .811
CH a(CH 2hCH2OH CH aCH 2CH(OH)CH a
48
-
~
Z
:I:
~ ~
o
§o Z
00
-
-
203*
16,030
14,740
-
216
16.150
14,860
-
-
-
-
209*
-
-
210*
-
-
9.31
-
-
-
344
8250
7340
CH aCH 2C(OH) (CHah
88.1
215.8
15
.815
6.79
-
(CHahCHCH2CH t OH
88.1
269.2 -179
.814
6.78
-
(CHahCHCH(OmCH.
88.1
233
.825
6.87
(CH a)aCCH20H
88.1
236
-
62.1
387.5
GLYCOLS AND GLYCEROL Ethanediol-l,2 (Ethylene Glycol) ..................... CH2(OH)CH 2OH
-
6.63 464.0 78.7 0.272
32.0
120-125
9
1.118
..
Propanediol-1,2 (Propylene Glycol) ................. CH3CH(OH)CH 2OH Propanediol-1,3 (Trimethy- CH 2 (OH)CH2CH 2(OH) lene Glycol)
76.1 76.1
371 850 (appr .)
Propanetriol-1,2,3 (Glycerol). CH2(OH)CH(OH)CH 2OH
92.1
554
ETHERS Methyl Ether .............. CH,OCH 3 Ethyl Ether ............... CH,CH 2OCH 2CH,
46.1
-
94.1
-
-
273* 266*
-
-
-
260
52
5.99 381
35
8.68
65.0 1.265
10.53
-11.5 -217
74.1
-
1.042 -
-
-
-177 .3 0.719
-
-
-
10,350 10,450
9350 9450
-
7760
6940
0.271
187
13,570u
12,340u
.262
151
15,840
14,560
-
129 120
16,930 16,870
15,630 15,570
17,560
16,250
-
-
Propyl Ether .............. CH 3(CH 2)20(CH 2)2CH 3 Isopropyl Ether ........ '.' .. (CH')2CHOCH(CH,)2
102.2 102.2
194.2 -188 155.3 -122
.752 .729
6.26 6.07
-
Butyl Ether ............... CH 3(CH 2)30(CH2)3CH 3 Sec. Butyl Ether ........... [CH3CH 2CH (CH,) 120
130.2 130.2
288.0 -144 250 -
.773 .760
6.44 6.33
-
-
-
115* 109*
-
-
-
-
-
320*
.786
6.54
-
-
-
257*
11,400
10,540
ALDEHYDES Methanal (Formaldehyde) ... HCHO
30.0 -
3
-180
-
-
8050U
-
7420U
~
P::
~ 00
......
(1
> ~
Ethanal (Acetaldehyde) ..... CHaCHO
44.0
Propanal lPropionaldehyde)
58.1
120
-114
.812
6.76
-
-
-
215*'
13,400
12,420
72.1
167.2 -144
.809
6.74
-
-
-
189*
14,640
13,590
00
72.1
142
.799
6.65
-
-
-
180*
14,600
13,550
58.1
133.0 -138.8
.795
6.62
-
-
-
220
13,260
12,280
> Z ;1
72.1
175.5 -123.5
.810
6.74
-
-
-
190
14,540
13,490
86.1 86.1
216.1 -108.0 215.2 - 40
.812 .820
6.76 6.83
-
15,430 15,380
14,330 14,280
200.7 -134
.820
6.83
-
168* 168*
86.1
-
165*
15,350
14,250
100.2
240.6 -119
.806
6.71
-
-
152*
15,980
14,840
CH,CH2CHO
Butanal (Butyraldehyde) .... CH 3CH 2CH2CHO 2-Methylpropanal (Isobutyraldehyde) ............... (CH 3)2CHCHO KETONES Propanone (Acetone) ....... CH 3COCH a Butanone (Methyl Ethyl Ketone) ................. CH 3COCH 2CH, Pentanone-2 (Methyl Propyl Ketone) ................. CH 3COCH 2CH 2CH, Pentanone-3 (Diethyl Ketone) (CH 3CH 2hCO 3-Methylbutanone-2 (Methyl Isopropyl Ketone) ........ CH 3COCH (CH a) 2 4-Methyl Pentanone-2 (Methyl Isobutyl Ketone) .
CH,COCH~H(CH,h
68.5 -190.3
-
-
-
87
-
-
* Calculated or estimated with a probable accuracy of ±2%.
(1
o
~
~
"'-J
u Heat of combustion as a gas-otherwise as a liquid.
..
PHYSICAL CONSTANTS OF GASES
FOn~ULA
MOLEC. WT.
BOILING POINT
MELTING POINT
of
of
CRITICAL CONSTANTS
HEAT OF COMBUSTION
@
60°F-BTU j1b
,~~
P Atm
D G /ml
Gross
270.3
111.5
0.235
9670
8000
88.0
73.0
.460
-
-
-
-220.4
34.5
.301
4345
4345
291
76
.57
-
-
7.51
369
51.6
.33
-
-
-400
12.8
.031
61,100
51,600
124.5
81.6
.42
-
-
°API
Sp Gr Lb/gal 60°/60°
97.5
0.617
5.15
'42.0
.815
6.78
t of
NH 3
17.0
-
Carbon Dioxirle ......
COz
44.0
-109.3*
-
Carbon Monoxirle ....
CO
28.0
-312.7
-::137.0
-
-
Chlorine .............
Ch
70.9
-
30
-151
-
-
Ethyl Chloride .......
CzH~CI
64.5
54.1
-214
Hydrogen ...........
Hz
2.0
-423.0
-434.5
-
-
-
Hydrogen Chloride ...
HCl
36.5
-121.0
-173.6
-
-
-
Hydrogen Sulfide .....
H 2S
34.1
-
76.5
-122.0
46.0
.797
6.64
212.7
88.9
-
7100
6550
Methyl Chloride .....
CH 3Cl
50.5
-
11.6
-143.8
20.3
.931
7.76
289.6
65.8
.37
-
-
Nitrogen ............
N2
28.0
-320.5
-346.0
-
-
-
-232.8
33.5
.31
-
-
Oxygen .............
Oz
32.0
-297.4
-362.0
-
-
-
-181. 9
49.7
.43
-
-
Sulfur Dioxide .......
S02
64.1
14.0
-
-
315.0
77.7
.52
-
-
Ammonia ............
28.1
-107.9
DENSITY
69.9
98.9
25.5
.901
1.394 11.62
"'tl ~
~ ~
o > ~ o o
z
U1
"-3
>
Z
"-3
U1
*.Sublimes.
~
..
Section 2
CHARACTERISTICS OF PETROLEUM FRACTIONS Average Boiling Point of Petroleum Fractions
,.
Many physical properties of pure hydrocarbons can be correlated with specific gravity and normal boiling point as independent variables. However, for use in the petroleum industry, these correlations must also be applicable to petroleum fractions which are mixtures of a large number of components, usually having a wide variation in boiling points. While the average specific gravity is a property of the petroleum fraction which can be measured directly, just as in the case of pure compounds, there is not an analogous average normal boiling point for a mixture. By integrating or averaging its distiUation curve (temperature vs. liquid volume percent distilled), a volume average boiling point can be determined for the mixture. However, as Watson and Nelson! and Smith and Watson 2 have pointed out, this has no special significance as a true average boiling point and many physical properties can be better correlated by the use of some other average boiling point, i.e., weight average, molal average, etc. Consequently, in all correlations involving boiling points of petroleum fractions, the proper average should be used. For the following physical properties, these are: Average Boiling Point
1 2
Physical Property
Volume average
Viscosity Liquid specific heat
Weight average
True critical temperature
Molal average
Pseudo-critical temperature Thermal expansion of liquids
Mean average
Molecular weight Characterization factor Specific gravity Pseudo-critical pressure Heat of combustion
Watson and Nelson, Ind. Eng. Chem. 26, 880 (1933). Smith and Watson, Ind. Eng. Chem. 29,1408 (1937). 10
CHARACTERISTICS OF PE'l'ROLEUM FRACTIONS
11
Since a distillation curve is usually available and a volume average boiling point is readily obtained therefrom, the other average boiling points are given as a function of these data. The chart on page 14 is based on an assay (True Boiling Point) distillation 3 of the whole crude, while the chart on page 15 refers to the 1070 (or ASTM) distillation of the fraction itself. The chart on page 14 was derived empirically from crude assay fractions of a number of crudes. For narrow boiling fractions, all of the average boiling points approach each other and the volume average boiling point may be used for any of the others. Then, by appropriately combining the volume average boiling points of the narrow cuts, the various average boiling points of wider cuts were determined. The weight and molal average boiling points of the wider cuts were calculated directly by combining the narrow cuts on the basis of their weight and mole fractions, respectively. The mean average boiling point could not be calculated in the same manner since it is not a direct average or integral of its fractional parts. As used herein, mean average boiling point is defined as the boiling point which best correlates the molecular weight of petroleum fractions. Consequently, the mean average boiling point for wider cuts was determined indirectly from the generalized molecular weight chart on page 21. Although Smith and Watson proposed a cubic average boiling point for the correlation of characterization factor, specific gravity-boiling point relations forthe different crudes indicate that the present mean average boiling point can be used for correlating gravity, and consequently characterization factor. Smith and Watson also used cubic average boiling point for correlating viscosity, but the present data indicate that the volume average is the proper boiling point. Since these boiling point correlations were developed directly from crude assay distillations, this chart should always be used 4 if an assay is available. Otherwise, the 10% (or ASTM) distillation of the fraction may be used in conjunction with the other chart. The latter was derived from the crude assay chart and an empirical correlation between the two types of distillation curves. The difference between the two sets of curves at zero slope represents the thermometer stem corrections for the 10% distillations. In the case of light hydrocarbon mixtures, where the analysis is known, the volume, weight, and molal average boiling points can be calculated directly from the boiling points of the components and their volume, weight, and mole fractions, respectively. On the oth~r hand, the mean average boiling point must be determined indirectly from the average molecular weight of the mixture. Up to an Approximately 15 theoretical plates and 5 to 1 reflux ratio. Below slopes of 2°F/% for low boiling fractions (V.A.B.P. < 500°F) and 3°F/% for high boiling fractions (V.A.B.P. > 500°F), the volume average may be used for the other average boiling points with very little error. 3 4
DATA BOOK ON HYDROCARBONS
12
average molecular weight of 80, the molecular weight-boiling point relation for normal paraffins (page 20) may be used for this purpose, but for higher molecular weights the generalized chart on page 21 should be employed. Characterization Factor Watson and Nelson 1 introduced characterization factor as an index of the chemical character of pure hydrocarbons and petroleum fractions. The characterization factor of a hydrocarbon is defined as the cube root of its absolute boiling point in oR divided by its specific gravity (60°F/60°F), or Characterization Factor
=
yt T B/Sp Gr
Characterization factor is given on page 16 as a function of gravity in °API and boiling point in of for hydrocarbons and petroleum fractions. That characterization factor is only an approximate index of the chemical nature of hydrocarbons is indicated by its variation with boiling point both for members of a homologous series and for fractions from the same crude (page 17). However, it has considerable value in that it can be applied to the entire boiling range of a crude and it has been generally accepted by the petroleum industry. Typical Crude Fractions For approximate use when there are insufficient data, several correlations have been developed for typical crude fractions grouped according to characterization factor and viscosity index. 5 These groups are numbered in order of decreasing paraffinicity and each may be considered representative of the crude fractions within its characterization factor or viscosity index range. The five groups were arbitrarily selected as follows:
Group I II III
IV
v
Characterization Factor . 12.1-12.6 . 11.9-12.2 . 11.7-12.0 . 11.5-11.8 . 11.3-11. 6
Viscosity Index of Lube Fraclions6 80-100 60-80 40-60 20-40 0-20
Fractions from some of the more common crudes are cla5sifil'd in the following table: lS
6
See page 156. Dewaxed to +20°F pour.
CHARACTERISTICS OF PETROLEUM FRACTIONS
13
TYP.lCAL GROUP
White Products Pennsylvania I Rodessa. . . . . . . . . . . . . . . . . . . . . . . . . . . .. I Panhandle . . . . . . . . . . . . . . . . . . . . . . . . . . . II Mid-Continent . . . . . . . . . . . . . . . . . . . . . .. II Kuwait I-II CRUDE
Gas Oils and Heavier I I I II II-III
Iraq Iranian East Texas South Louisiana. . . . . . . . . . . . . . . . . . . . .. Jusepin
II II III III III
II-III II-III II II III
West Texas . . . . . . . . . . . . . . . . . . . . . . . . .. Tia Juana (Med. and 102) Colombian Lagunillas . . . . . . . . . . . . . . . . . . . . . . . . . ..
III III IV V
III IV IV V
Since, in the case of some crudes, the lower boiling fractions belonged in a different group than the higher boiling fractions, they were classified separatelythat is, into white prorlucts having an average boiling point less than 500°F, and gas oils and heavier having an average boiling point greater than 500°F.
+40
+ 30
2
3
4
5
6
WEIGHT ·AVERAGE·
+20 +'0
0 - 10
7
-
10
•
Ii
iJ
9
~
AVERAGE BOILING POINT OF PETROLEUM FRACTIONS ~ CRUDE ASSAy DISTILLATION
- 20
i
MEAN AVERAGE
-30 -40 -50 -60 -70
2
4
0
5
6
7
8
THE SLOPE AND THE 50% POINT FOR THE VOL. AV. B.P. UNLESS THE DISTILLATION FOR THE FRACTION DEVIATES APPRECIABLY FROM A STRAIGHT LINE. !N THE LATTER EVENT THE FOLLOWING FORMUI AS SHOULD BE USED:
-40 -
_ t7O-t10 S 60 ty = to+4t.50+tIOO 6
-60 In
..
cr....
('~
10
* THE CUT RANGE MAY BE USED FOR
MOLAL AVERAGE
-20
V>
9
-80
FOR WHOLE CRUDES:
t y = ho
-'00 -120 -'40
3
4
5
7
6
14
8
9
10
t
t,50+ teo
t40 WEIGHT AVERAGE +20 ,.L
0
-20
,~ 4
2
5
t . t·
6
7
8 i:J...~t-L_; of
MEAN AVERAGE
i....... .ly.......H-t~.
AVERAGE BOILING POINT
+20
,.
OF PETROLEUM FRACTIONS 0
10 % (A.5.tM.) DISTILLATION
-20 IF AVAILABLE, THE CRUDE ASSAY DISTILLATION SHOULD BE USED FOR DETERMINING AVERAGE BOILING POINTS.
-40
r r-
l/}
2). Vaughan and Graves, Ind. Eng. Chem. 32, 12.;>2 (1940). Wiebe and Brcevoort, J. Am. Chem. Soc. 52, 622 (1930). Wiebe, Hubbard and Breevoort, J. Am. Chem. Soc. 52, 611 (1930) .
•
_~:
.-- -
4-
::l
HH-
'!' --
80 70 60
50
40 -_.~
30
20
-;-
-300
o
-200
27
100
200
-2'50
-200
-150
.8
-
.1 .6 .5
l--:r.
~
--
-100
~
.. - , ~ ~.-
VAPOR PRESSURE OF ETHANE AND ETHYLENE
.4 .3
.2
200
.08
80
.07 .06
60
.()5
50
70
.04
40
.-
.O~
30
.02
20
all', , II
.1 ~
_.
...-
gr
4 3
2
-200
-tOO
o
100
28
200
300
400
-100 - ,-;;.,,",
i-
-50
::T:-E.=:io€-~:ff-'
~ ::ri..:r--,~f.':'=j-:',
le"",
-,
..6
,,,="
=
-'
.1
'~7,k
,c.:.~3':!..J
,
= VAPOR PRESSURE OF PROPANE AND PROPYLENE
.5 .4
+- ,
'+:
.3 --H
.2
.1 .09 .08 .07
200
100
~
r=
90
_
'''''
" !CA_~{~!
.06
-t-t=f
60
-,.!~
.05
70
''f'
50
.04
40
.03
30
.02
20
II -'I=.:f.: "
.009 ,008
-'
.001
10 9 8 7
IXXJ
6
.005
5
=r
4
T
3
2
o
100
200
300
29
400
500
50
·5
-
.4
VAPOR PRESSURE OF BUTANES AND 8UTENES
:II+-_~,
..3
.2
200
30
I .
II II II
.01 ~-====47
-
1
I I
/I
I
o
V V 1/
I
tOO
200
300
30
400
500
-roo
t.O .9
o
50
p.
: ::r=tfel. ·-:-r-·: ~-.g:
--
VAPOR PRESSURE OF .5 ~
PENTANE
~
-I'
...
::T'
AND ISOPENTANE
.3
.2
'200
11
I
.€ ==':.f"_ -=I::
:t=:!-~:.
90
_ !:::f
.07 .06
1=l;.T:,:
80 70 60
.05
50
.()4
40 30 20
.02
.OOQ
Q'
I
I
.01
.- -
, .
-',--.
-:.
,008 .007 .006
.:
-
-
:~
±
.§'
- ":'::. - .
--
8
7 6
,=:
l:±-
H
.005
5
.004
4
.003
_.
- I::±l
2
.002
It
.
.00IU-J....l....L.LLL.JLJ...,U.I-......AJ...J...Iu1t.U./-l..'~~I~J.:..z.J,;.W-l...l-J...J.-l...J..J..1..J..J...J...L..J...J..J..J..J...J...IW-L...J.-l....J.-l...J..J..u..J-U...J...J..J..J..1..J..J...J...L-U 100
200
300
400
500
31
l--.
.-.---------
600
-50
-100
0
100
- 1 I f t i j - -.:.~
.9
.8
..,
';-..::
T.~:
.
':
="=±
.5
-
50
: .....
'.':~:
ISO
VAPOR PRESSURE OF HEXANE
.4
.a
200 I~
.1 .09
-t...
.
-fl
:.~,"'f> :. 'c: :::'.-:x':~ ::~-,.
-17
.07
.
n-
"
_.
"c't._',_ - _"':4'_ .•••.-.t
'--J-;,'
:: t~
:'':'
--. :r.' ,-- '-,-, , 'f--X:==, T.:i::-'~ ~,
.-
..
-'-+-
.
~.
=- __.---7l~:. ==1':::.
,.. -,-V-FFl r-- ~-;:=::;=:I -I:' ' =l:
'
, -
100 90
80 70
.06
60
.05
50
.04
40
.03
30
.02
20 .'
r
I
/
;=;=. ~ 11.,,':.. '.,~':'1.
£
.008 .OOf .006
.'f=:£:i=E j~ --!'
.''=8, !E::-l,=.:l3:: ' --:=;::1:'. :t+= 'r:
- -,
1-..
-- -
,~.
,~f:~T ~.~:~I _:. ,_ ,. ~ -
,-J.'
~t:.:~
'
-, -. . ,.. =.-1-=!iiC:l 10 9
= -
9 7 6
5
.00
4
3
.003
2
I
I I I
I
II .• ;.,
32
-50 .8
.5
0
. =-=
::::.C
50
100
150
·=--==~~-:'~';'_~c~"=§=:l~F-=~. .
_--
...
.4.0 0 F/'Y.
0 COORDINATING RESEARCH
COUNCIL (CRC) HANDBOOK. PP. 244-254 (1946)
44
Q.
~ L&J
Q.
0
0..
0:
It
It
0:
10
REFERENCE:
400
0
TRUE VAPOR PRESSURES CORRESPONDING TO A R.V.P. OF 9.0 LBS./ so. IN. AND A SLOPE OF 4.0°F/'>'. ARE READ FROM THE CHART AS FOLLOWS: TEMPERATURE TRUE VAPOR PRESS. OF LBs/sa. IN. 32 2.8 100 9.9 21.4 150
10
It (/)
l&J
SLOPEc 160-120
20
300 w ;:) (/)
EXAMPLE: DETERMINE THE TRUE VAPOR PRESSURES AT 32°F. 100°F AND 15Q°F OF THE FOLLOWING GASOLINE: REID VAPOR PRESS. - 9.0 LBS./SO. IN. DISTILLATION + LOSS, 5'>'. ~ 120 OF 15'>'. (! 160 OF
30
~
:::>
12 14 16 18 20
40
~
(J) (J)
II
50
0
~ en 5.0 CD
4
;;: 4. •0 5 2 6 UJ 7 Q:" 8 :> en 9 en 10
L&J 0:
80
~ d
...J
l&.
y
*
SLOPE
120
90
50 60
160
100
40
II
12 13 14 15 16 17 18 19 20
30
500 ~ us ;:) 600 It I-
700 800 900 1000
Section 5
FUGACITY Raoult's Law If two or more compounds form an ideal solution in the liquid phase, and if the saturated vapors of the individual components are perfect gases, the system has been termed an ideal system.! For such a system the partial vapor pressure of any component may be calculated from the composition of the liquid phase by Raoult's Law and from the composition of the vapor phase by Dalton's Law. An equation of these two expressions gives the liquid-vapor equilibrium relation for any component, i where i = 1, 2, ... , n:
or
Pi = PiXi = 7rYi
(1 )
== P i /7r = Ki
(2)
yi/Xi
where Pi
=
partial pressure of i
Pi = saturated vapor pressure of i Xi = mole fraction of i in the liquid phase
Yi
= mole fraction
7r = Ki
of i in the vapor phase
total (vapor) pressure of the system
= vapor-liquid equilibrium constant for i at the temperature and pressure of the system
The above equation, usually referred to as the Raoult's Law relation, is true only for ideal systems, ,as defined above. However, it is usually a good approximation for mixtures of homologues and, in general, for mixtures of chemically similar compounds, if none of the saturated 'vapors at the equilibrium temperature deviate too greatly from a perfect gas. Up to moderate pressures (several atmospheres) hydrocarbon mixtures frequently fall within the scope of the Raoult's Law relation. However, its application to these mixtures is rather limited because of the wide differences usually encountered between the boiling points of the most volatile and least volatile components. This results in equilibrium temperatures at which the saturated vapors of the lowest boiling components deviate considerably from a perfect gas, even though the equilibrium pressure of the system may be relatively low. I
Gamson and Watson, Nat. Petroleum News} Technical Section 36, R-258 (1944).
45
DATA BOOK ON HYDROCARBONS
46
Fugacity Functions In order to improve the accuracy in predicting vapor-liquid cquilibrium constants for hydrocarbons at higher pressures, Lewis and Luke 2 and other investigators replaced the pressures in equations (1) and (2) by -analogous fugacities for any component, i, whereby:
or
Ii = fpiXi = fnYi
(3)
Ki
(4)
Yi/Xi
where fi
=
=
fpdf.Tri
=
fugaci·ty of i in either phase of the system
fpi = fugacity of i as a pure saturated liquid (or vapor) at its vapor pressure
corresponding to the equilibrium temperature of the system f-rri
= fugacity of i as a pure vapor at the equilibrium temperature and pressure of the system
Generalized correlations have been developed for the r-atio of fugacity to pressure for pure hydrocarbons as a function of reduced temperature and reduced pressure. A correlation of this type (pages 62 and 63) was used in conjunction with the vapor pressure charts to develop the fugacity function charts for individual hydrocarbons. 3 The fugacity function given by these charts, 7rfp/!-rr, may be considered a corrected vapor pressure and used in place of the latter in ~ny equation pertaining to liquid-vapor equiiibrium such as equations (1) and (2). These simple fugacity relations greatly extend the pressure range for which liquid-vapor equilibria for hydrocarbon systems may be predicted with confidence, and can be used up to equilibrium pressures of 20 to 25 atm with a fair degree of accuracy. Beyond these pressures and especially as the critical point of the mixture is approached, serious deviations from true equilibrium conditions are encountered. Under these circumstances, the assumptions of ideal mixtures no longer hold and the fugacities of the individual compounds are dependent upon the compositions of the liquid and vapor phases as well as temperature and pressure. In the region where the simple fugacity relations no longer apply and consequently beyond the scope of the present charts, there are data in the literature on a number of specific binary and multicomponent hydrocarbon systems. Also, The M. W. Kellogg 00. 4 has published an excellent correlation for lig~1t paraffin and olefin hydrocarbons in which the fugacities of the individual compounds are given as a function of the molal average boiling points of the liquid and vapor 2 Lewis
and Luke, Trans. Am. Soc. M echo Engrs. 54, 55 (1932).
a This method was actually used only up to the critical temperature of each compound. Beyond this point values were calculated from more general fugacity correlations developed by The M. W. Kellogg Co. to avoid using extrapolated vapor pressure curves. 4"Liquid-Vapor Equilibria in Mixtures of Light Hydrocarbons," The M. W. Kellogg Co., New York, N. Y. (1950).
FUGA CITY
47
phases in additi on to the equilib rium tempe rature and pressu re. The Kellogg correl ation was derive d from the applic ation of exact therm odyna mic relatio ns to a comprehensive equati on of state for pure hydro carbon vapors and liquids and their mixtur es. 5 If, in additi on to hydro carbon v-apors, other gases (air, H , CO , etc.) are 2 presen t in the vapor phase, it is recom mende d that an effective pressu 2 re, equal to the produ ct of the total pressu re multip lied by the square root of the mole fractio n of the entire hydro carbon portio n of the vapor, or 7r-VV;;;, be used in determ ining the fugacities of the indivi dual hydroc arbons . Fragm entary data have indioa ted that this effective pressu re gives better results than either the total pressure, 7r, or partia l hydro carbon pressure, 7r'YHC, for determ ining indivi dual fugacities. Then, after the fugacities or fugaci ty functio ns have been read from the charts , the total pressure is again used as a basis for all equilibl:ium calcul ations . The following examp le illustr ates the applic ation of the fugaci ty functi on charts when other gases are presen t in the vapor phase: Examp le 1. Determ ine the pressu re and composition of the liquid phase in equilib rium with a vapor of the following composition at gO°F: 1st Trial Component
Vapor. Mole Fract.
1l'
0.040 .220 .280 .175 .160 .125 1.000
CH 4 C2H6 C3Hs C 4H 1O
= 25 atm
1I"e = 21.5 atm F, atm
Air H2
2nd Trial
* * 180 38.0 13.5 4.9
x
-
0.039 .115 .296 .637 1.087
Interpo lation
11" = 20 I1tm 1I"e = 17.2 atm
F, atm
11"
x
* *
-
180 36.0 12.7 4.4
0.031 .097 .252 .[;68 0.948
= 21.8 atm x
0.034 .104 .269 .593 1.000
• In this example, the fugacity functions of air and H2 are conside red to be infinite.
where 7r = total equilib rium pressu re 7re = 7rVO.740 = effective pressu re used to determ ine fugaci ty functio ns F = 7rfp/!1I" = fugaci ty functi on for pure hydroc arbons
x ~
= 7rY/ F
Relati ve Volati lity Since relativ e volati lity is quite useful in fractio nation problems, curves for the relativ e volatil ities of light unsatu rates and isoparaffins to the corresponding norma l paraffins are given on pages 64 to 66. The curves for the C 4 unsatu rates IS
Benedict, Webb and Rubin, J. Chem. Phys. 8, 334 (1940); 10, 7474 (1942).
48
DATA BOOK ON HYDROCARBONS
inay also be used in conjunction with the normal butane fugacity chart to predict fugacity functions for these compounds. Except for butadiene and the normal butenes, these relative volatility curves were derived from the Kellogg fugacity correlation. Composition was indirectly taken into account to some extent since the fugacities for each pair of compounds were read at the same liquid and vapor molal average boiling points as well as at the same temperatures and pressures. In general, the relative volatility charts may be considered to have a somewhat greater range of applicability than the simple fugacity charts. They may be used up to 25 atm, irrespective of the composition of the liquid and vapor phases of the mixture; beyond this pressure their application is limited to systems in which there is a difference of at least 75°F between the molal average boiling points of the two phases, but under no circumstancel? should the curves be extrapolated. While all of the curves may be considered to be accurate within 25% for the relative volatility minus one (0; - 1), deviations from the solid curves rarely exceed 15% for this difference. Chemical Structure and Liquid Activity Coefficients When components in a hydrocarbon mixture are quite dissimilar chemically, the liquid phase may deviate appreciably from an ideal solution. This effect of chemical structure is not taken into account in any of the fugacity correlations heretofore considered. It has been mentioned that in correlations of the Kellogg type, fugacity is a function of the liquid and vapor compositions, but only with respect to components of similar chemical structure. To correct for chemical dissimilarity in solutions of light hydrocarbons in absorber oils, liquid activity coefficients are given for these light hydrocarbons on page 67. Within the range of the data these activity coefficients were practically independent of temperature (100°F and 220°F) and pressure (500 psia and 1000 psia) .
GENERAL REFERENCES Brown, Souders and Smith, Ind. Eng. Chem. 24, 513 (1932). Dean and Tooke, Ind. Eng. Chem. 38, 389 (1946). Hadden, Chem. Eng. Progress 44, 37 (1948). Kay, Chem. Revs. 29, 501 (1941). Lewis, Ind. Eng. Chem. 28, 257 (1936). Lewis and Kay, Oil and Gas J. 32, 40 (1934). . Lewis and Randall, "Thermodynamics," pp. 190-198, McGraw-Hill Book Co. (1923). Nelson and Bonnell, Ind. Eng. Chem. 36, 204 (1943). Sage and Lacey, Ind. Eng. Chem. 30, 1296 (1938).
~
..
·~-t-toj-or
~T:~
FUGACITY FUNCTION
1m
OF METHANE
_. t
-l- .
200
r v.
'I i . ':.;=
~:-:.:F-~:-~-
-;':~i:£
-.
J:-
10
-::-±
_
I
.• -~
100
90
I::f.;=+
80 70
60
50 40 30
20
-
I
: :.-;;!
~:.,~
=~~~ -'- :.=;-
j. p-L~-t.: :~- .
_
~/J::.1~·_-,2? ~E§= -
•
.j...
-
=:=.
10 9
8 7
6
5 ,
4
..::~
3
-300
-200
o
-100
49
100
200
300
&I .
..
F~~~CITY FUN~TI~N OF ETHYLENE
I
r +-
-
..
I-- ~
1 00
90
80 70
60 50 40 30
20
-
••
_I~ 8 7 6
5 4
-200
o
-100 50
100
200
300
~
-~=
~
-
--- --
-
-
---
=
= = = = = -
FUGACITY FUNCTION OF ETHANE --
-
~
00
•
--
--
-
-
-
--
-
- - --
-
--
-
-
-
- ~t a -
--
I 00
-- -
-
90
---
-
80
-
70 60
--
50
40 30
20
~
-
-
~
/ -
--
- -rl=Fl
10
-'
- -
9
8 7 6
5 4
3
2
-zoo
o
-100
51
100
zoo
300
1.0 0.9 0.8 0.7
0.6
--200
o
-100
.=a=::-
-
/M
0.5
.
-
-c
FUGACITY FUNCTION OF PROPYLENE
0.4
-
-
0.3
0.2 I~
/1
AI If I
II
0.1
I 00
90 80 70 60
...
50
40 30
20 I-
-, I I
IJ
iii
II
10 9 8
7 6 5 4 3
2
-100
o
100
52
200
300
400
-200
1.0 0.9
0.8 0.7 0.6
-
o
-100
, 0:;
FUGACITY FUNCTION. OF PROPANE ::
i
0.5
- . -.
0.4
-
-+-+-l
-
0.3
0.2
..
0.1
~
I 00
i
90 80 70 60 50
40
-
30
20
11 -(
If
V I
iW
II
'.
•
, •
I
8
7 6
4
~
l.
-
9
5
-+-
-l-
10
3
.
,
,
2
I
"'T
I
J
I
i,""'" I
-ir-
,
I
I
!..ll -j(jO
__.x. ~
L
If
.-L~
c.:.'_ _,-
o
100
53
I I
200
300
400
__
:T"-~
,- FUGACITY FUNCTIO:lI "OF ISOBUTANE :: ,
~
-
r
t
0.2
"1J;:,f-;t--l+++H Ifl+I-I-/HI-H+-I--+ ,
I
'I
I
ElIOO
~l--
~
=3...J .
~l='-
90
80 70 60
,-
50 40 ___ t:::t:::t:=t=J:
30
20
I.-
t;o
~
r+- -H-lH+I--t+H+t++-H+-H+-l-H-++++H+H-I-+-1-++I-I4-IA+:I4-jll--W-+-I-l--!-l-l--+-W-!-W~-I--l 1
I ~:-:.!-
3:,-'=i-- -i-:-(=F" -...t.~ ~-- .~~.l
'7
--
""--
It
=-
-
-~Ig
..;
--t.-
8 7 6 5 4
3 -I- -
1/111 1/
I
-tOO
o
tOO
54
200
300
400
o
100
-!!III
·i-tll- f:.Tfl ~aii--~f~ "':::-l_
~---.-.
--r=..t=
:.
-~
8+
-,
I-
-f-
FUGACITY FUNCTION [~I--
r:
§i=-I
OF BUTANE
- ffP-i' -:-.:~. '- Ii § ~~ H-".IEI:- ~=8::=
_,.-,_-r-F
--=+:~-
-
--'-+-
'--4-=1= 0.2
1111
I
--.
100 90 80 70 60
j:-t-;'
±:'--; -I--l.-r·~·
7
.. -f~- 3f~
,.J~ 51~ =t--+=!'I !_-l ~~:
- ..: ->-.-
t-
50
,~
f---
...::;C :-I.,
~
40
i7
30
.-
20
1/ j
=. -.
-,..
---:::-:~-_.
~r_ ~_r=-- ~ ~ -~_
1=I j:::f
...
, .-1'\)
10 9 8 7 6
-f-
5 ..
.
. ..•
~E:::
r-
4
-
:;=. :7t;- ~ 3
I I
I
H--i--H-+-t-l---H-++-+-+-+-IH-+++-t--HH-++-+-+--HH- 'I - - LIl-lH--H..j-+-HH--H-++-HH--H-++-H-+-H-++-H-+-~--H-j I
I
-100
o
100
55
200
300
400
-100
a
100
0.9 0.6 0.7
~
0.6
0.5
~.
FUGACITY FUNCTION OF ISOPENTANE ~ :
0.4
0.3
0.2
t I I
40
30
20
I I
4 3
I
IjjjT±-;1-±±ttttjjjttt::t:t:tttttlltJzt~~:tt:t1tttt:!i!~~~~:t;;'tr~t'ttt:r~;r~11!!!t:!=1
o
100
200
56
300
400
500
1 H-++++ -H-++-- H-t-f-+- t--H 't+t-i-:+t+'
1
., -"1f-f-+I H--H-+ t-H-t-f- +-H-+t +t-H-t- f-+-H-+ t+t-Hf- f-t-I-t-t --H-H- Hf-f-t --:
50
=
40
... .
17'
•
1
I~
:¥T 1
17'
20
If
11
It
30
, .+-H-H-H-H-i+H-H-H-H-H-H-H-H-H-+++++
. I
-
+++++-hI"'H-T+-H'-H' '.s the latent heat of vaporization of benzene at 1 atm. Example 2. Determine the latent heat of vaporization of the following ga9 oil at 500°F. 10% Distillatioll 10% @ 430°F 50% @ 540°F 70% @ 605°F 90% @ 680°F
Gravity 35°API
Vol. Av. B.P. = 547°F; Slope = 2.9°F/% Mean Av. B.P. = 547 - 9 = 538°F Molec. wt. = 211 Vapor pressure (538°F normal B.P.) = 0.63 atm. at 500°F Pseudo-critical pressure = 266 psia co 18.1 atm I'l'lolee. wt. of normal paraffin (538°F normal B.P.) = 222 Critical pressure of normal paraffin = 15.0 atm Vapor pressure of normal paraffin = (l5.0/18.1)0.63 = 0.52 !!.tm Latent heat of normal paraffin = 104 BTU/lb Latent heat of vaporization of the gll3 oil!!.t 500°F =
22Z X 104 = 108 B1'Ulib 214
78
DATA BOOK ON HYDROCARBONS
Enthalpy of Light Hydrocarbons The enthal py 8 or heat content of low-boiling paraffins, olefins, and aromatics is given by the chart.s on pages 98 to 113. These charts can be applied to mixtures of light hydrocarbons on the basis of the following assumptions:
1. The entha!pies of individual components of a mixture are additive in the !iquid phase, that is, the mola! heat content of the mixtttre equals the sum of the products of the mo!al heat contents of the components by their mo!e fractions. 2. The entha!pies of individua! components are additive in the vapor phase at !ow pressures (0-1 atm). 3. The change in enthalpy of the vapor with pressure at constant temperature is the same for a mixture as for a sing!e compound having the same mo!ecular weight as the mixture. The first assumption is substantially true for hydrocarbon mixtures (especially for homologous series) at temperatures below the critical regions of all components. At temperatures near to or above the critical temperatures of any of the components, the liquid mixture is no longer an ideal solution of its components and there is some deviation from the rule of additive heat contents. However, since these deviations arc not too serious, and since no other simple method has been developed for determining the heat content of a liquid mixture, the rule of additive enthalpies should be used for all hydrocarbon mixtures irrespective of the critical temperatures and chemical composition of the components. The second assumption is strictly true only for vapor mixtures at infinite dilution (0 atm) but is a very close approximation for pressures up to 1 atm. The third assumption is empirical but has been shown indirectly to give quite accurate rcsults for mixtures of homologous series and petroleum fractions. Also, the usc of the average molecular weight to determine the change of enthalpy with pressure is the simplest average which can be used. Above thc critical temperature a dashed line is shown for the heat content of the gas in solution. This line was based on the assumption that thc gas in solution at any temperature would have the same partial density and enthalpy as the pure compound at a pressure corresponding to an extrapolation of its vapor pressure curve above the critical point. Obviously, this is only a rough approximation since both a vapor prcssure curve and an ideal liquid solution arc meaningless in this regIOn.
E:rample 3. Determine the difference in enthalpy bctween the liquid at 100°F and the vapor at iiOO°F and 20 atm for a mixture having the following composition: 8
Based on on entholpy of zero for the saturated liquid
01,
-200"F.
THERMAL PROPERTIES Component
Mole Fraclioll
C,H. CsH. C,H lO C,H, CsH.
0.100 .500 .100 .050 .250 1.000
79
The enthalpy of the mixture as a liquid at 100°F and as a vapor at 500°F and 0-1 atm is computed from the individual components as tabulated below:
Com-
Mole
poncnt
Fract.
CzH s
0.100 .500 .\00 .050 .250
CaH s C 4 H 1O
C2H 4
C3 H,
Enthalpy of Liquid lOO°F
l\lolcr, Wt.
Enthalpy of Vapor 5OQ°F and 0-1 elm
Lb/Mole of Mixture
BTU 1~lole
BTU lIb
3.0 22.0 5.8 1.4 10.5 42.7
of Mixture
239 171 159 223 169
li y (500°F, 0-1 atm) - li L
BTU lIb
720 3760 920 310 1770 7480
BTU ll\lolc
of J\'lixture
553 530 525 506 508
1660 11660 3040 710 5330 22400
= 22,400 - 7480 = 14,920 BTU/mole
The change of enthalpy of the vapor at 500°F betwecn 0-1 atm. and 20 atm. is computed by interpolating between C,H. and CsH.: C,H.: 1f,,(500°F, 20 atm) - li y (500°F, 0-1 atm) = 30(546 - 553) = -210 BTU/mole CsH.: H y (5OO°F, 20 atm) - li y (500°F, 0-1 atm)
=
44(522 - 530)
=
-350 BTU/mole
Mixture: li y (500°F, 20 atm) - li y (500°F, 0-1 atm) = -210 =
- 30 + 42.7 H _ 30 {-350 -
(-21O)J
-340 BTU/mole
Therefore,
li y (500°F, 20 atm) - liL(100°F) = -340
or
14,580 42.7
= 849
+ 14,920 =
14,580 BTU/mole
BTU/Ib
The foregoing procedure can hc simplificd, with a loss of accuracy which does not usually exceed 5%, by interpolating on a basis of molecular wcight and total
so
DATA BOOK ON HYDROCARBONS
olefin content between the initial and final states:
CaH s: H v (500°F, 20 atm) - Ih(IOO°F) = 44(522 - 171) = 15,440 BTU/mole
C ZH 4 : H v (500°F, 20 atm) - Ih(lOO°F) = 28(500 - 223) = 7750 BTU/mole
CaHe:
H v (500°F, 20 atm) - HL(lOO°F) = 42(500 - 169) = 13,900 BTU/mole
Since the average molecular weight of the paraffin portion of the mixture is 44, the propane values can be used directly, making interpolation unnecessary. The average molecular weight of the olefin portion is 39.7; hence the enthalpy difference between the initial and final states will be: 7750
+ 3:;7 _- 2~8 (13,900 -
7750) = 12,880 BTU/mole
Interpolating between the paraffin and olefin portions, Hv(500°F, 20 atm) - HL(lOO°F)
or
X 15,440 + 0.30 X 12,880 = 14,670 BTU/mole
= 0.70
14,670 42.7 = 344 BTU/lb vs, 342 BTU /lb by the longer method.
Enthalpy of Petroleum Fractions The enthalpyO of petroleum fractions is given by the charts on pages 114 to 127 for both paraffinic stocks, having a characterization factor of 12.0, and nonparaffinic stocks, having a characterization factor of 11.0 over a mean average boiling point range from 200°F to 800°F. Theoretically, these charts represent pure hydrocarbons of the designated characterization factor and boiling point, but they may be applied to petroleum fractions if the following assumption is made in addition to the three previous ones pertaining to light hydrocarbon mixtures: 4. 'The avemge difference between the enthalpy of the vapor at low preSSU1'es (0-1 atm) and the enthalpy of the liquid, at constant temperature, is the same f01' a rni.'l:ture of chemically similar hyd1'Ocarbons as for a single compound of the sct'1ne molecular weight (or mean avemge boiling point). \\'hile this assumption is empirical, it is accurate \\'ithin a few percent excrpt in the region of the pseudo-critical temprrature \\'here the enthalpy of the liquid is subject to variation depending upon the true criticaltemperatUl'e of the mixtUl'e. Since the dashed line starting at the pseudo-critical point applies only to a pure compound in solution above its critical point, another dashed line was arbitrarily drawn for mixtUl'es, joining the satUl'aled liquid line below the pseudo-critical 9 Based on all enthalpy of zero for the saturated liquid at 00 F .
•
THERMAL PROPERTIES I
81
point with the pure compound line about 50°F above the pseudo-critical tempera_ ture. This is more representative of a mixture and should be used in preference to the pure compound line. These charts may be interpolated and extrapolated linearly with both characterization factor and mean average boiling point. Occasionally, in inter·
polating between two adjacent boiling point chart~ the pressure and temperature of the vapor will be such that they fall inside of the "dome" of the higher boiling point chart.. Since it is impossible to use the charts in this region, it is recommended that the two adjacent lower boiling point char'" be extrapolated upward. Following are two examples illustrating the use of these charts: Exan,ple 4. Determine the ditTcrence in enthalpy between the liquid at 500°F and the vapor at 775°F and 25 psig for the following refined oil fraetion: Grauity 40 0 API
Crude Assay DistillatiOlt I.RP. 300°F 50% 440°F F.B.I'. 580°F Vol. Av. RP. = 440°F
· '11 . 81 ope 0 f t}Ie dIstl atlOn eurve
=
580 - 300 100
=
2.8 0 F/%
Mean Av. B.P. = 440 - 6 = 434°F Characterization Factor = 11.65 hI' = Enthalpy of the vapor at 775°F and 2.7 atm (25 psig) h L = Enthalpy of the liquid at 500°F Mean Ch. Cb. Cb.
Au. B.P. -400°F Factor = 12: h" - hL = 567 - 286 = 281 BTU/lb Factor ~ II: h v :- h L = 538 - 263 ~ 275 BTU/lb Faetor = 11.65: hI' - h L = 275 + 0.65(281 - 275) = 279 BTU/lb
Mean Ch. Ch. Ch.
Au. B.P. -500°F Faetor = 12: hI' - h L = 556 - 273 = 283 BTU/lb Factor = II: hI' - h L ~ 534 - 255 ~ 279 BTU/lb Faetor = 11.65: hI' - hi = 279 + 0.65(283 - 279) ~ 282 BTU/lb
Mean Au. B.P. - 434°F Cb. Factor = 11.65: hI' - hL = 279
+ 1.'.. (282
- 279) =.280 BTU/lb
If the char'" for 300°F and 400°F Mean Av. B.P.'s had been extrapolated, the result would have been essentially the same, 281 BTU/lb.
Example 5. Determine the difference in enthalpy between the liquid at 425°F and the vapor at 925'F and 350 psig for the following gas oil:
82
DATA BOOK ON HYDROCARBONS 10% 10% 50% 70% 90%
VoI. Av. B.P. = 455
Distillation @ 455°F @ 560°F @ 620°F @ 695°F
Gravity
15.5°API
+ 2 X 560 + 695 = 5670F 4
. SI ope 0 f di Bt I'II ation curve
- 455 = 620 69 = 2.8°F/% 5 = 562°F
Mean Av. B.P. = 567 Characterization Factor = 10.48
h v = Enthalpy of the vapor at 925°F and 24.8 atm (350 psig) h L = Enthalpy of the liquid at 425°F Av. B.P.-400°F Factor = 12: hv - h L = 662 Factor = 11: h v - h L = 622 Factor = 10.48: h v - h L = 406 Mean Av. B.P.-500°F Ch. Factor = 12: hv - hL = 642 Ch. Factor = 11: h v - lt L = 606 Ch. Factor = 10.48: hv - h L = 398 Mean Av. B.P.-562°F Ch. Factor = 10.48: ltv - h L = 388 Mean Ch. Ch. Ch.
233 = 429 BTU/lb 216 = 406 BTU/lb 0.52(429 - 406) = 394 BTU/lb 224 = 418 BTU/lb 208 = 398 BTU/Ib 0.52(418 - 398) = 388 BTU/lb M(394 - 388) = 385 BTU/Ih
MollieI' Diagrams The MollieI' diagrams for the individual light hydrocarbons are of essentially the same type as the familiar one for steam. To minimize confusion and to make the charts as easily usable as possible, lines of constant volume are omitted and lines of constant temperature replace lines of constant superheat in the superheated vapor region. These charts will be used principally for adiabatic compressions and expansions. In applying the MollieI' diagrams to hydrocarbon mixtures, the mixture should be treated as a single compound of the average molecular weight. An empirical study of the diagrams indicates that successive charts of the same series (paraffin or olefin) may be interpolated (or extrapolated) by assuming a linear relation exists between melecular weight and (1) isentropic change of molal enthalpy with pressure and (2) the product of the square root of the molecular weight and the isentropic ehange of temperature with pressure. If both paraffins and olefins are present in the mixture, the charts of each
•
,
THERMAL PROPERTIES
83
I
series are interpolated (or extrapolated) to the average molecular weight of t.he total mixture. These values corresponding, respectively, to a 100% paraffin mixture and a 100% olefin mixture are used for linear interpolation to the actual olefin content of !lie mixture. The following example illustrates the application of the MollieI' diagrams to a hydrocarbon mixture: Exmnple 6. Determine the work of compression 1 0 and final temperature when
the following mixture is compressed adiabatically from atmospheric pressure and 60°F to 50 psig: Average Molee. WI.
M ok Fraction
Component
CH, . ....... . C,H, ........ C2 H• ....... . C3 H. C3 H s ..... - .. C,H s C.H, 0 C.H, .......
0.050 .100 .150 .100 .200 .100 .200 .100 -1000
,
0.8 2.8 4.5 4.2 8.8 5.6 11.6 7.2 -45.5
Values corresponding to adiabatic compression from 1 aIm and 60°F to 4.4 atm were read from the individual charls and arc tabulated below: BTU/lb Compound
C,B, C~HIO
C2H.. C3H e
0.763 .680 .935 .780
By interpolation,
lif{ =
6fl M(h, - h,)
I,
S h,
112
301.5 295 300.5 303
338 321 363.5 3!2
'F 1M 135
1610 1510 1760 16iO
221
164
1600 BTU/mole and litVM
61VM
BTU/mole
=
625
570 850 675
620 for a saturated hydro-
carbon mixture of 45.5 malec. wt.
By extrapolation, li.H ~ 1610 BTU/mole and li/V,lf hydrocarbon mixture of 45.5 molec. wI.
=
632 for an unsaturated
By interpolation, li.H ~ 1603 BTU/mole and litVM
=
6z.t for a hydrocarbon
mixture of 45.5 malec. wt. containing 30% unsaturates. 10 Change in enthalpy which includes the difference between the work of expulsion and work of admission.
DATA BOOK ON HYDROCARBONS
84
:. The theoretical work of compression is 35.2 BTU/lb and the final tempera.ture is 152°F. If other gases (H 2 , O2 , H 2 0, etc.) are present in a mixture, it is recommended that effective IJressures equal to 7rVYlfC be used to determine the total work of compression and final temperature of the hydrocarbon portion of the mixture. The inert g&ses usually may be assumed to be ideal and the w'ork of comprestlion and final temperature for this portion of the mixture calculated by the adiabatic compression formulas for perfect gases. The work of compression for the mixture is then evaluated by combining the change of heat content for the hydrocarbon portion with that for the inert gases on the basis of their mole fractions. In determining the final temperature of the mixture, it is assumed that the ch.ange in enthalpy of each portion from its final temperature to that of the mixture is equal and opposite in sign to the other. This method is illustrated by the following example: Example 7. Determine the work of compression and final temperature when
the following mixture is compressed adiabatically from 25 psig and 0°1" to 150 psig: Hydrocarbon Portion Average Molec. Wt.
Mole Fraction
Component
Average Molec. Wt.
-
-
1.0 1.6 4.5 11.0
0.500 .100
H, CH, C,H. C,H.
Mole Fraction
.ISO
.250 1.000
0.200 .300 .500 1.000
18.1
3.2 9.0 22.0 34.2
The effective pressuras to be used for the hydrocarbon portion of the mixture are: 7rei
=
7r.2
=
25
+ 14.7 14.7
150
vO.500 = 1.91 atm
+ 14.7 vO.500 -=
14.7
7.91 atm
Values read from the ethane and propane charts are tabulated below: BTU/lb Compound
S
C,H. CaR.
0.837 0.686
t,
h,
h,
294 278
340 307
6.H
OF
M(h, - hi)
MvM
121 92
1385 1280
663 610
THERMAL PROPERTIES
85
By interpolation, UTi = l35~ BTU/mole and t:.tVM = 6-1.7 for a saturated hydrocarbon mixture of 3-1.2 molee. wt. The corresponding final temperature for the hydrocarbon portion of the mixture is 111°F. For the H 2 portion of the mixture, the work of compression and final
tcm.pera~
ture arc calculated as follows:
MCp
=
UTi
=
2.016 X 3.46
~
K
6.96;
-.
=
6.97 99 6.97 - I.
1.40
~I R1'[("2)K~' - IJ
J( -
"1
[(1647)1.40-' I J
I. 40 X 1.99 X 460 __ . 1.40 1.40 - I 39.7 = 3200(1.502 - I) = 1610 BTU/mole =
I.fO -1
7'2 = (164.7)1:'lE: !4+ a.. ~ "« Hill' I 'H. l. ~ ';iit Z
It
l'
I'
t-!! I '
~
~l..
f
t ~!.
••
I"'"
t
;1'
j
'11"_
~
•
-I
I
, 1
••
1."
••••
Ih·
-.
~
'j
:.;..
, ......,
.; "" '"
; l i .•~'~ • •
I
....
.'!~
"..!~~
''t'
;',1
. ..• - I: .".
~;:; Ii .. '''' ...
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';
Htill ill "I+H: lj:Hl++hr; tii:nr !~ri ;LJ \ :: '-:~ ..!i;! jl~l L1':': :!;j ~;:l £::1 I';q :~!; ,I;1.n.. t; f. II r~ '; ~\' .~\., d . -.-.. • ;t; . ,.It.tl-.l n:j I,; .. , .pl ;'\,.; . .'~~~ .Li~'. I,U .. 1. ;,.~ I. ... ,t. r" 'jl ,~'. W ' 't. !lfl;' I'll r .II!' r 111't '1. li1 :t;" t,'~~n;,-,,! 1'1 '"" ,.W .'
350
56
340
s.,
330
300
5,3
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~
290
~
280
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270 260 250
100
55
320 310
110
59
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>iii
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230
'"
70
8
..J
!!! >
0
'">-
80
w
0
8w
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90
60
-
in 0
tJ:;
UI
....
(j'
80
f--
t
z
850
~
....
f-!:::
T
2300+ 2200t 2100
f T
2000
0
0'" ":;'"
'" :;
70
60
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al
>ow .-(/) -ct
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8.,
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129 177 273 369 465 561 656 752 848.
-7
0.40
-
.36 .34 .34 .34 .33
-11
-19 -22 -23 -20
-
-
.33
+57
Flash Vaporization (F) t>t'
FRL
Curve
-3
233 265 329 393 457 521 585 649 713
230 265 325 387 450 513 578 649 .732
-
-4 -6 -7 -8 -7 -
+19
The flash reference line and the atmospheric flash curve of the original crude are-shown on Figure 1. Proceeding from (1), the flash curve of the original crude, the atmospheric flash curves of the stripped and unstripped reduced crudes are derived by the method outlined in the text: (2) Vol. Av. B.P. of whole crude
=
262
+ 538 + 905 = 568°F 3
226
DATA BOOK ON HYDROCARBONS Mean Av. B.P. of whole crude = 568 - 70 = 498°F (Section 2) Molec. wt. of whole crude = 197 (Section 3) Vol. Av. B.P. of 65% overhead = Slope of DRL (65% overhead) =
203
+ 373 + 3
558
= 378°F
495 - 139 60 = 5.9°F/%
Mean Av. B.P. = 378 - 38 = 340°F Molec. wt. of 65% overhead = 139
(3)
(4)
(5)
(6)
Per 100 Gal of Crude Moles of crude = (6.98 X 100)/197 = 3.55 Moles of overhead = (6.57 X 65)/139 = 3.07 Moles of reduced crude 0.48 Partial pressure of reduced crude at the dew point of the original crude 0.48 = X 1 = 0.135 atm. 3.55 The 40% point on the reduced crucle flash curve corresponds to 65 + 0.40 X 35 = 79% or 722°F on the flash curve of the original crude. By extrapolation from 0.135 atm. to 1 atm., the 40% point on the atmospheric flash curve of the reduced crude is 900?F. The atmospheric flash curve of the stripped reduced crude is drawn through the extrapolated point parallel to the 65-100% portion of the flash curve of the original crude. This reduced crude flash curve may be converted to percent on reduced crude by proportioning the 65-100% yield on original crude to 0-100% on reduced crude. Both curves are shown in Figure l. The front end of the atmospheric flash curve on the unstripped reduced crude is constructed by drawing a smooth curve from the 65% point on the flash curve of the odginal crude to the 20% point on the flash curve of the stripped reduced crude as shown in Figure 1. This curve is also given on the basis of 0-100% reduced crude.
GENERAL REFERENCES Edmister and Pollock, Chem. E7l{J. Progress 44, 905 (1948). Katz and Brown, Ind. E7l{J. Chern. 26, 1373 (1933). Packie, Trans. Am. Inst. Chern. Engrs. 37, 51 (1941).
EQUILIBRIUM FLASH VAPORIZATION
227
1100
1000
900
800
700
600
!lOO
400
300
200
10
20
30
40
50 FIGURE
60 1
70
80
90
100
7
PREDICTION OF FLASH REFERENCE LINE FROM DISTILLATION REFERENCE LINES
--JfJ1 .•r
~
5
.. ~.
7
g
w
4
2
. FLASH AND DISTILLATION REFERENCE LINES (FRL AND DRLl ARE STRAIGHT LI NES THROUGH THE 10% AND 70% , POINTS. THE TEMPERATURES AT THE
I -
o
i li
If
3
50% POINTS REFER TO THESE REFERENCE LINES.
3
6
7
8
9
10
II
,,'
2
m I .
12
.~ I~
40
~ ~
E : ... RE~G~l~it!fI:~:':I~-'~ 1
.... ;~~ 2
3
o
60
t
•
T
i ~If
U4
6
S
7
!'iii R,f;f,Et EN' ~...
!; - .
8
9
10
-
II
12
11- -
40
20
..
II
BI
O -20
-40
-60
-40
2
3
5
6
228
tl±I
.
~
.
, ,
~
,
.
;
, tmI,
I
: t 1.
..
,
Mm c
1
,
, •
too
1.00 .80
lm!l
100
I..
PREDICTION OF FLASH CURVE FROM ITS REFERENCE LINE
I
•
1'-'-
fmII
mIl tmn
,
II
11UlI
•
_ 11$ CRUDE , .
.60
ASSAY (T.B.P.) DIS TlLLATION
'gmj
,20
11m!
-,,",p .. I:J±
t"
10
,
1.00
20
30
II
40
50
11m
60
70
90
100
.
·lmJU
.JI#J:
"* lIV IS THE DEPARTURE
n5
.60 f:fk,'ll
I
OF THE ACTUAL FLASH AND D1STI LLATION CURVES FROM THEIR RESPECTIVE REFERENCE LINES. WHILE THE INDIVIDUAL (lIl')'S MAY BE EITHER PLUS OR MINUS, THE RATIO IS ALWAYS POSITIVE.
,
~ , ' '. ..
~l'mIJ
Iflill
.20
o
80
lEE ' 10% (A.s:tM.) DISTILLATION
.80
!l0
•• • II
10
W
30
·w_ •
40
~
60
ro
229
~
90
100
.1
Section 14
FRACTIONATING TOWERS In order to simplify the work involved in making stepwise calculations for the rectification of binary and multieomponent systems, Gilliland' has presented an empirical correlation between theoretical steps and reflux ratio. To use the Gilliland correlation to predict the number of theoretical plates for a given reflux ratio, the minimum number of steps at total reflex and the minimum reflux ratio are required. Minimum Number of Theoretical Steps When a separation is specified with respect to only two components of a multicomponent mixture, the lower boiling of these two components is designated the light key component and the higher boiling the heavy key component, and the minimum number of steps can be calculated by the well-known Fenske equation 2 as follows: a log
8 Af=
(XXLKlV LKD) (X HKlV) XHKD log aLK
[aLK]SM =
or
(XXLKlV LK D) (X HKlV) XHKV
(1)
(1 a)
After equation (1) is solved for 8M , the latter may be substituted in this equation along with the distribution of either key component to prcdict 4 the distribution of the other components, or
(X
,.D) X HKlV) = 8 M log aL XX log ( -
(2)
LK log (X lilY) (X LKD) = 8 M log (a ) XlID XLKlV all
(3)
!-IV
11 K D
Likewise,
In any of the above equations, moles per 100 moles of feed may be replaced by total mo)es, or volume or weight units since in any of these conversions the multiplying factors cancel out. Gilliland, Ind. Eng. Chern. 32, 1220 (1940). Fenske, Ind. Eng. Chern. 24, 482 (1932). A table o( nomenclnture is given on page 243. • This equation may be used (or any pair o( component8.
1
2 S
230
FRACTIONATING TOWERS
2H1
When the dcgree of separation is specified for more than two components, equation (1) must be applied to all critical combinations of these components and the maximum SJ/ determincd for the most difficult case. If the separation is specificd with respect to the total quantity of two or more components, as in the case of Examplc 1, trial and eITor is required for thc solution of SjJ. It should be pointed out that the concentrations calculated by equations (2) and (3) actually apply only to the separation at total rcflux and, with the
exception of the two key components, there will be some variation of thc degrec of separation with the reflux ratio. As the rcflux ratio decreases, there is some improvement in separation betwecn light and heavy componcnts boiling outside the range of the kcy components and some deterioration in the separation of components boiling intermediate bctween the kcy components. However, in so far as the present procedure is concerned, the distillate and bottoms compositions for other reflux ratios are assumed to be thc Same as those calculated for total reflux. Minimum Reflux Ratio
Gilliland 5 has proposed several diffcrent formulas for predicting minimum reflux ratio and all have the disadvantage of being composcd of a number of complex terms in addition to requiring trial and error for solution. Although all these equations appeal' to give satisfactory rcsults, the tcrms are so complcx that it is difficult to bc ccrtain that therc arc no numerical crrors in thcir application. In order to apply the Gilliland method with greater facility, the following equation was developcd for predicting the minimum rcflux ratio of a multicomponent system: (O/Dhf
+ 1=
(aLKTf.FC + ('(LK - 1
1) (XLKD ILK
XIIKD)
(4) (O/D)M can be calculated for two arbitrary states of feed vaporization: 1. "Liquid" feed, cOITesponding to vaporization of the feeu equivalcnt to the
fraction of the feed lighter than the light key component. For the components lighter than the light key, h = ZL/aL and for the light key and heavier components, ILK = ZLK, and III = Zfl.6 Gilliland, Ind. Eng. ehe",. 32, HOI (1940). ]£ components, intermediate between the two key components, are present, they are considered ei her light or heavy componen'" depending upon which key their volatility more nearly approaches. In the case of "liquid" feed, I L = Zl. and I II = ZH for these intermediate components; in the case of "vapor" feed, I L = ZL/aL and I II = ZuaH/aLK. 5
8
232
DATA BOOK ON HYDROCARBONS
2. "Vapor" feed, corresponding to vaporization of the feed equivalent to the fraction of the fecd consisting of the hcavy key component and lighter. For the components lighter than the heavy key, If_ = ZL/CtL and ILK = ZLK/aLK and for the components heavier than the heavy key, 1/1 = ZI/.6 After the minimum reflux ratios have been calculated for "liquid" and "vapor" feeds, the minimum rcflux ratio for the actual vaporization of the feed can be calculated by direct interpolation or extrapolation. However, extrapolation beyond 50% of the difference between "liquid" and "vapor" feed may lead to serious deviations. The first term of the right-hand side of equation (4) is the same as for binary mixtures, and the equation reduces to the cquivalcnt of a binary mixture whon all light components other than the light key have infinite volatility and all heavy component other than the heavy key have zero volatility. Under these circumstances the equation is exact when hK is taken as the ratio of the two components in the liquid phase of the feed. That is, if the feed is introduced as a liquid at its bubble point, hK = ZLK, which is the ratio of the two components in the feed; if the feed is introduced as a vapor at its dewpoint, hK = Z',K/OtLK, which is the ratio of the two components in the cquilibrium liquid. For intermediate stages of vaporization hK can be calculated from the flash vaporization formula, although direct intcrpolation of the minimum reflux ratio on the basis of percentage vaporization between thc sat urated liquid and saturated vapor feeds gives values only slightly in error on the conservative side. In the case of multicomponent mixturcs, equation (4) is semi-empirical since it was necessary to make simplifying approximations in its derivation. Furthermore, the exact values of the various 1's cannot be calculated directly from the composition and state of vaporization of the feed, since the liquid on the feed platc is not identical to the liquid phase of the fecd as in the case of a binary mixture. As a result, it was necessary to define the 1's empirically for two states of fced vaporization, arbitrarily choscn to simulate a binary mixture of the two key components, and then intcrpolatc or extrapolatc to the minimum reflux ratio corresponding to the actual vaporization of the fced. Equation (4j has been checked for a number of multicomponent systems on which the minimum reflux ratio was determined by stepwise trial and error calculations. Generally, unusual systems were chosen with respect to composition and relative volatility in order to reveal the maximum deviations ever likely to be encountercd in practice. The agrecment was quite satisfactory as the average deviation was less than -+-5% and the maximum about lOra. The latter occurred at the limit of extrapolation relative to the arbitrary feed states.
•
FRACTIONATING TOWERS
233
Also, the minimum reflux ratio was calculated for these same systems by the Colburn method 7 with about the same degree of accuracy. It should be pointed out that the latter gave better results than equation (4) when the relative volatilities and compositions were not so abnormal as the systems selected. However, under these circumstances both methods were quite accurate as the deviations seldom exceeded a few percent, and the present equation has a distinct advantage in that it is explicit and does not require trial and error. Both methods are quite sensitive to the selection of key components, and the selection of the wrong key components can lead to a much greater error than is inherent in either method. If the desired separation is between adjacent components, there is usually no doubt about selecting these as the key components. However, if there are additional specifications relative to other components, it may be necessary to try two or more combinations of key components to make sure that the minimum reflux ratio is sufficient to fulfill all specified conditions. Correlation of Theoretical Steps with Reflux Ratio As mentioned at the beginning of this section, Gilliland correlated the results of a large humber of stepwise calculations on various binary and multicomponent mixtures by plotting [S-SM) I[ S + 1] ~ (S) against [(OlD) - (OID))f)/[ OlD + 1)-F (OlD)
and found that all points could be represented by a single curve irrespective of the type or degree of separation. These points, along with about half again as many additional points, were replotted, and the best curve through them was essentially the same as Gilliland's original correlation. In arriving at the coordinates for the additional points the minimum reflux ratio was calculated by equation (4); therefore these points are a criterion of the present method as well as the curve itself. In no case did the deviations exceed either 3 theoretical steps or 15%, and the average deviation was less than 1 theoretical step and also less than +50/0. To take care of the maximum deviation it is recommended that in any design the number of theoretical steps predicted ~ from the correlation on page 244 be increased by either 3 theoretical steps or 100/0, whichever is greater. Plate Efficiency Because of the large number of factors which undoubtedly influence the plate efficiency of a fractionating tower, any fundamental formula accounting for even the most important variables must necessarily be quite involved. For this reason, a simple empirical correlation of the limited data on hydrocarbon mixtures seemed to be the most promising method of predicting plate efficiency. 7
Colburn, Trans. Am. Inst. ehem. Engrs. 37, 805 (1941).
DATA BOOK ON HYDROCARBONS
234
Gunness 8 correlated the results of several tests on petroleum mixtures on the basTs of vapor pressure of the liquid. As he points out, this is a method of indirectly correlating plate efficiency with liquid viscosity since viscosity of pure hydrocarbons and narrow boiling fractions is an approximate function of vapor pressure over a fai,rly wide range of vapor pressures. In view of the consistent results obtained by Gunness, pla.te efficiency was plotted directly against fluidity (reciprocal viscosity) for a number of tests on commercial towers including those upon which Gunness based his curve. The curve on page 245 represents this correlation. While the overall plate efficiency exceeds 10070 at fluidities greater than 9 Cp-1, this is not inconsistent as the flow of the liquid across the plates results in concentration gradients which may achieve a greater degree of fractionation than predicted by stepwise calculations in which the liquid is assumed to leave the plate in equilibrium with the composite vapor. Lewis 9 has shown theoretically that different combinations of liquid and vapor concentration gradients across the plate may give overall plate efficiencies as high as 200--300'10 when based on stepwise calculations. There is no reason to believe that this correlation applies to mixtures other than hydrocarbons, and with the exception of alcohol-water mixtures there are too little data available to afford a comparison. Although there is considerable variation in the alcohol-water data, there is some indication that plate efficiencies are somewhat greater than for hydrocarbons of the same viscosity. Location of the Feed Plate As a simple approximation for locating the feed plate, it may be assumed that the proportion of actual plates above the feed will be the same as that required to effect the same separation between the key components at total reflux. That is, the number of theoretical steps at total reflux is calculated for the concentration change in the key components between the feed and distillate compositions. It is then assumed that the ratio of this to the total number of theoretical steps at an infinite reflux ratio is the same as the ratio of actual plates above the feed is to the total number of plates. Application of this method is illustrated by Example 1. In some cases where there are oritical components other than the two key components, it may. be necessary to check the total reflux steps above and below the feed on the basis of these components, since the optimum location of the feed plate will be different with each pair of components. Usually the separation of components other than the key components is so complete that only the latter need be considered. 8 9
Gunness, Sc.D. Thesis, Mass. Inst. Tech. (1936). Lewis, Ind. Eng. Chern. 28, 399 (1936).
FRACTIONATING TOWERS
235
Packed Towers The charts on pages 246 to 248 giving the H.E.T.P., capacity and pressure drop in packed towers are self-explanatory. Since practically all of the H.E.T.P. data were on towers less than 12 in. in diameter, caution should be used in the design of larger towers. One of the greatest sources of inefficiency in a packed tower is poor liquid distribution. If good distribution can be achieved by efficient distributors, the extrapolations may be used for larger towers with reasonable assurance. Example 1. At an operating pressure of 100 psig determine the number of plates and reflux ratio required to separate the mixture given below so that the bottoms contain at least 90ro of the butenes-2 present in the feed and at the same time have an isobutene content not greater than 5%:
Component
i-C,H lo i-C,H s C,H s-1
C.H IO t-C,H s-2 c-C,H s-2
Feed (Mole %) 40.0 20.0 15.0 5.0 10.0 10.0 100.0
(1) Dewpoint of Distillate and Bubble Point of Bottoms
In order to calculate the average volatilities, the dewpoint of the distillate and bubble point of the bottoms must be found by trial and error using assumed compositions. These are tabulated below.
Moles Per 100 Moles of Feed Component i-C.H LO i-C.H, C.H...1 C.H,o t-C.H,·2 .,.C.H...2
Mole Fraction
Feeti
Distillate
Bottoms
Distillate
Bottoms
40.0 20.0 15.0 5.0 10.0 10.0 100.0
39.3 18.7 13.0 1.0 1.5 0.5 74.0
0.7 1.3 2.0 4.0 8.5 9.5 26.0
0.530 .253 .176 .014
0.027 .050 .077 .154 .327 .365 1.000
.OW .007 1.000
As a first trial, assume the dewpoint of the distillate is 14tl°F at 7.8 atm (114.7 psia) .
DATA BOOK ON HYDROCARBONS
236
First Trial Component
YD
':"C,H IO i-C,H, C,H,-I C,H,o t-C,H...2 c-C,H ...2
a'D·
Pt
:i;
HO°F
140°F
"yiP
1.29 1.155 1.13 1.00 0.97 0.91
8.4 7.5 7.35 6.5 6.3 5.9
0.493 .263 .187 .017 .025 .009 0.994
0.530 .253 .176 .014 .020 .007 1.000
• Relative volatilities to C 4H 10 or (0")'8 aTC used as a matter of convenience; then, the (a'.,.)'s are converted to (a••)'s, the relative volatilities to t-C,H,-2, which will be seleetcd as the heavy key component. t Computed from the fugacity function of butane multiplied by the relative volatilities.
Since the sum of the x's is 0.994 instead of 1.000, the assumed temperature should be lowered slightly, but the difference would be so small (less than l°F) that the change in the (a'Jl) 's would be imperceptible. Consequently, 140°F will be used as the dewpoint of the distillate. The bubble point of the bottoms is assumed to be 165°F at 8.0 atm 10 for the first trial.
Second Trial
First Trial
Component
i-C,H,o i-C,H, C,H ... I C,H,o t-C,H ...2 c-C,H,-2
XIV
0.027 .050 .077 .154 .327 .365 1.000
a' w· 165°F
pt
Y
165°F
Pxl..
1.26 1.14 1.115 1.00 0.97 0.915
10.7 9.7 9.5 8.5 8.25 7.8
0.036 .061 .091 .164 .337 .356 1.045
a'w· 160°F
160°F
Y
1.265 1.145 1.12 1.00 0.97 0.915
10.25 9.3 9.1 8.1 7.85 7.4
0.035 .058 .088 .156 .321 .338 0.996
Pt
• Relative volatilities to C~HIO or {«')'8 are used ns a mnttcr of convenience; then, the (a'av)'S are converted to (aU\')'s, the relative volatilities to t-C IH:;-2, which will be selected as
the heavy key component. t Computed from the fugacity funet:on of hutane multiplied by the relative volatilities.
The bubble point of the bottoms wiil be taken as 160°F. The relativ~ volatilities are averaged and converted to t-C~ H s -2 as the heavy key in the following table: \0
After allowing 3 Ib/sq in. as the approximate pressure drop through the tower.
FRACTIONATING TOWERS
,
,
,
aD
aw
aA
Component
140°F 7.8 at.m
160°F 8.0 atm
150°F 7.9 atm
i-C,H ID i-C,H. C,H.-l C,H ID t-C,H8-2 c-C,Hg-2
1.29 1.155 1.13 1.00 0.97 0.91
1.265 1.145 1.12 1.00 0.97 0.915
1.275 1.15 1.125 1.00 0.97 0.91
237
,
a a.
(a' Da'wa'.A.)~!l
1.275 1.15 1.125 1.00 0.97 0.91
aa.
1.315 1.185 1.16 1.03 1.00 0.94
(2) Minimum Theoretical Steps (Total Reflu:t) The minimum number of theoretical steps by which the desired separation can be accomplished is calculated as follows:
Let
t = moles of t-C 4 H s-2 in the distillate per 100 moles of feed 10 - t = moles of t-C 4 H s-2 in the bottoms per 100 moles of feed
Since 90% of the butenes-2 must be.retained in the bottoms, the cis-butcne-2 content of the distillate and bottoms will be: (2 - t) moles in the distillate per 100 moles of feed (8 + t) moles in the bottoms per 100 moles of feed
and
Using the previously assumed values of 18.7 moles of isobutene in the distillate and 1.3 moles in the bottoms, the following equations must be satisfied:
C1~;) CO t- t) 18.7) (~) ( 1.3 2- t
= (1.185)8 M =
(1.185)8 0.94
M
A trial and error solution of these equations shows that they are satisfied by SM = 25.5 and t = 1.62. The distribution of the other components can be calculated from SM and the distribution of t-C 4 H s -2. i-C 4 H IO : Let u = moles of i-C 4 H IO in bottoms (
40 u
1') (8.38) = (1.315)25.5 = 1075 1.62
= 0.19 moles of i-C4 H IO in the bottoms C 4H s-1: Let v = moles of C4 H s-1 in the bottoms (
15 v
v) (838) 1.62
=
(1.16)25.5
=
44
v = 1.58 moles of C 4 H s-2 in the bottoms
DATA BOOK ON HYDROCARBONS
238
C.H IO : Let w = moles of C.H IO in the bottoms (
5W
W)(8.38) = (1.03)25.5 = 2.12 1.62
W
= 3.55 moles of C.H 10 in the bottoms
The percentage of i-C.Hs in the bottoms will be: (0.19
+ 1.3 + 1.58 ~33.55 + 8.38 + 9.62) 100 = 5.3%
In order to meet a maximum of 5.0ro i-C 4 H s specified for the bottoms, it is necessary to reduce the 1.3 moles to 1.22 moles in the bottoms. This would require an increase in SM to 25.8 which would modify the distribution of the other components. However, the latter change is so slight that it can be neglected. The composition of the overhead and bottoms will then be: Mole Fraction
Moles Per 100 Moles of Feed Component
i·C,H,o i·C,H, C,H..1 C,H IO /·C,H,-2 .,.C,H..2
Fecd
Distillate
Bottoms
Distillate
Bottoms
40.0 20.0 15.0 5.0 10.0 10.0
39.81 18.78 13.42 1.45 1.62 0.38 75.46
0.19 1.22 ·1.58 3.55 8.38 9.62 24.54
0.528 .249 .178 .019 .021 .005 1.000
0.008 .050 .064 .145 .342 .391 1.000
(3) Minimum Reflux Ratio Since the critical separation is between isobutene and the butenes-2, the former is naturally selected as the light key component and trans-butene-2, since it is more volatile than the cis-butene-2, as the heavy key component. Butene-l is considered a light intermediate component because of the proximity of its relative volatility to that of isobutcne; normal butane is considered a heavy intermediate component since its relative volatility is nearer to the heavy key than the light key. The following tabulation gives the necessary information for calculating the minimum reflux ratios for the two arbitrary states of feed vaporization: Mole Fraction Component
':-C,H,o i-C,H, C,H ..1 C,H IO t-C,H,.2 .,.C,H..2
Type
L LK L H HK H
"'BV
Feed
Distillate
Bottoms
0.400 .200 .150 .050 .100 .100 1.000
0.528 .249 .178 .019 .021 .005 1.000
0.008 .050 .064 .145 .342 .391 1.000
1.315 1.185 1.16 1.03 1.00 0.94
HLiquid tl
lIVapor"
Feed
Feed
3.04 2.00 1.50 4.00
3.04 1.69 1.29 3.48
2.00
2.00
FRACTIONATING TOWERS
239
"Liquid" jeed-40% vaporized (O/D)
+1 = M
1.185 X 2.00 + 1.0 (0.249 _ 1.185 _ 1.0 2.00
+ 01. 31 5 (0.528 .315
1.03
(O/Dhf
° 1) . 02
3.04 X 0.021)
(0.249
+ 1.16 (0.178 0.16
)
0.94
+ 1.185 - 1.03 4.00 - 0.019 + 1.185 - 0.94 = 1.88 + 1.94 + 1.07 + 0.29 + 0.46 - 1 = 4.64
1.50 X 0.021)
(0.249 ) 2.00 - 0.005
"Vapor" jeed-90% vaporized (O/D) M
+1 =
1.185 X 1.69 + 1.0 (0.249 - - - 0.021 ) 1.69 1.185 - 1.0 1.16 ( + 1.94 + 0.16 0.178 1.03
+ 1.185 (O/D)M
1.03
1.29 X 0.021)
(0.249 ) 3.48 - 0.019
= 2.06 + 1.94 + 1.10 + 0.35 + 0.46
- 1
+ 0.46
= 4.91
Assume that the feed is sufficiently preheated to vaporize a percentage equivalent to the distillate or 75.4670. By interpolation, the minimum reflux ratio corresponding to this feed vaporization is: (O/D)M
75.46 - 40)
= 4.64 + ( 90 _ 40
(4.91 - 4.64) = 4.83
(4) Theoretical Steps vs. Reflux Ratio
Using the values determined in preceding sections for minimum theoretical steps,. 25.8, and for minimum reflux ratio, 4.83, the number of theoretical steps for various reflux ratios can be predicted from the correlation on page 244: OlD
F(OID)
4.83 5.25 5.75 6.50 7.50
0.067 .136 .223 .314
.
-
-
(8)
-
0.570 .502 .430 .366
-
8
Theoretical Platea-
61.3 52.7 46.0 41.3 25.8
60.3 51.7 45.0 40.3 24.8
.
.
- The reboiler i. considered the equivalent of one theoretical step. With a partial instead of a total condenser, a second theoretical step also could have been deducted.
DATA BOOK ON HYDROCARBONS
240
(5) Number of Actual Fractionating Plates
To predict the number of actual plates it is necessary to determine the average viscosity of the liquid on the plates. Since the temperature difference between the top and bottom of the tower is so small, the average viscosity may be taken as the viscosity at the average temperature. For this purpose the viscosity of butane at 150°F will be used. Viscosity of C.H IO @ 150°F = 0.216 cs "" 0.216 X 0.523 = 0.113 cp Fluidity = 1/0.113 = 8.9 Cp-l j Plate efficiency = 99% Using a plate efficiency of 99% the number of actual plates is computed for various reflux ratios: OlD 4.83 5.25 5.75 6.50 7.50
.,
S
Theoretical Steps
Actual Plates
.,
.,
.,
61.3 52.7 46.0 41.3 25.8
60.3 51.7 45.0 40.3 24.8
60.9 52.2 45.5 40.7 25.0
The number of actual plates is plotted against reflux ratio in Figure 1. A reflux ratio of 6.50 to 1, or 1.35 times the minimum, is selected. The number of actual plates corresponding to this reflux ratio is 45.5 so that a 50-plate tower would be required. (6) Location of the Feed Plate
The number of plates above the feed is based on the proportion of theoretical steps at total reflux which would be required to effect the change in concentration of the key components between the feed and distillate. This proportion is applied to the actual number of plates (including the reboiler) to determine the number above the feed plate. In order to take into account any appreciable difference in relative volatility above and below the feed, the relative volatility used for calculating the steps at total reflux between feed and distillate is the geometric mean of aD and a" or, an
=(
1.15)~i 1.155 0.97 X 0.97
=
1.19
The number of total reflux steps which would be required between the feed and distillate is calculated by the following equation: l8.78) (~) ( 20 1.62
=
1.19n
= 5.79' '
n = 10.1
FRACTIO ATING TOWERS
241
50
40
30
5
4
7
6 FIGUllE
8
1
Number of actual plates above the feed would then be: 10.1 (50 25.8
+ 1)
=
20
The vaporization of the feed can be taken into account by adding the fraction vaporized to n since 10010 vaporization would be equivalent to a theoretical step at total reflux. This would change the proportion of plates above the feed as follows: (
10.1 + I • 0.75) ( 00" +) 1 = 21.4 pates above the feed 20.8
Feed lines would probably be installed above the 2'!th, the 28th and 32nd plates from tile bottom of the tower.
242
DATA BOOK ON HYDROCARBONS GENERAL REFERENCES
Atkins and Franklin, Refiner Natural Gasoline Mfgr. (Jan. 1936). Brown, Sanders, Nyland and Hesler, Ind. Eng. Chem. 27, 383 (1935). Brown and Souders, Oil and Gas J. 31, 34 (1932). Chilton llnd Colburn, Trans. Am. Inst. Chern. Engrs. 26, 178 (1931). Elgin and Weiss, Ind. Eng. Chem. 31, 435 (1939). Fenske, Lawroski llnd Tongberg, Ind. Eng. Chern. 30, 227 (1938). Fenske, Unpublished data, Pennsylvania State College. Gilliland, Ind. Eng. Chent. 32, 918, 1101, 1220 (1940). Gunness, Ind. Eng. Chern. 29, 1092 (1937). Lewis and Wilde, Trans. Am. Inst. Chern. Engrs. 21, 99 (1928). Perry, "Chemical Engineers' Handbook," pp. 829-832, McGraw-Hill Book Co., New York, N.Y. (1941). Sherwood, Shipley and Holloway, Ind. Eng. Chem. 30, 765 (1938). White, Tram. Am. Imt. Chem. Engrs. 31, 390 (1935).
FRACTIONATING TOWERS
243
.vomenclature X x
moles of any component in distillate or bottoms per 100 moles of feed mole fraction of any component in liquid y mole fraction of any component in vapor D moles of distillate per 100 moles of feed o moles of reflux per 100 moles of feed OlD reflux ratio (OIDhl minimum reflux ratio corresponding to S = 00 S number of steps from still to distillate 8,1/ minimum number of steps corresponding to OlD = 00 P number of theoretical plates; with a partial reboiler and partial condenser, P = S - 2, and with a partial reboiler and total condenser,
P=S-l
ZH aD alV
LK HK L
H D W n m
ratio of mole fraction of any light component to heavy key component in the feed ratio of mole fraction of light key component to any heavy component in feed relative volatility of any component to heavy key at the dew point of the distillate relative volatility of any component at the bubble point of the bottoms relative volatility of any component at the arithmetic average temperature of the dew point of the distillate and the bubble point of the bottoms mean relative volatility of any component, (aD' alV . a.4)fi used as a subscript to refer to the light key component used as a subscript to refer to the heavy key component used as a subscript to refer to any light component used as a subscript to refer to any heavy component used as a subscript to refer to the distillate used as a subscript to refer to the bottoms used as a subscript to refer to the plates above the feed used as a subscript to refer to the plates below the feed
1IIIriCORRELATION OF THEORETICAL STEPS WITH REFLUX RATIO MULTICOMPONENT AND BINARY MIXTURES
.9
.8
.6
.5 .4 .3
.2 .1
244
OVERALL PLATE EFFICIENCY vs. FLUIDITY OF LIQUID ON PLATES
120
1.11111
120
110 100
100
90
90
80·
80
70
ONLY DATA 00 HYDROCARBON MIXTURES WERE USED IN THIS CffiRELATION, AND THERE WERE INSUFFICIENT DATA ON OTHER TYPES TO JUSTIFY A MORE GENERAL USE. HOWEVER, THERE WERE SOME EVIDENCE THAT THE CURVE IS A LITTLE CONSERVATIVE FOR ALCOHOL - WATER
60
40_ 50
70 60 50
MIXTURES.
40
30 20
: :• • ll!IIflE:IJffi 10
_ _ _ _ _ '0 2
3
4
5
6
7
245
8
9
10
II
12
13
14
2
3
45678910
20 .:_.
30 :t';
HEIGHT EQUIVALENT TO A THEORETICAL PLATE
4a!f11i1 2
(I) WHILE THIS CORRELATION WAS DE'
VELOPED fROM DATA ON RASHIG RINGS AND 8ERL SADDLES, IT PR08ABLY APPLIES TO OTHER SIMILAR TYPES OF HOLLOW PACKING. (2) VALUES OF H.E.lP. FROM THIS CHART CORRESPOND TO THE MAXIMUM TOWER CAPACITIES GIVEN BY THE CHART ON THE OPPOSITE PAGE. FOR THE VALUES OF HE.T.P. AT CAPACITIES BETWEEN 80% AND 100% OF THE MAXIMUM, DIVIDE H.E.T.P. FROM CURVES BY THE FRACTION OF ULTIMATE CAPACITY (.80-1.00) AT WHICH THE TOWER WILL OPERATE.
mIR~12ioll~3IoI14~oll~ 60
• 246
7'08090 I
·1
* USE VALUES OF S/F 3 FROM CUR\IE
~-d
FOR RASCHIG RINGS. BERL SADDLES W PACKING UP 10 2 INOiES IN SIZE. FOR SIZES GREATER'THAN 2 INCHES, USE INDIVIDUAL
.2
.,.,, ~li
VALUES OF SAND F.
.1
lilt! 1)
r r 1111
'r-
.08
~
jj
.06 .05 .04
.
e:l3
.f+
"t~
I 1 I
ltllllltltrTl1111 ;~':+fTo"",
1000 800
.
,I';r.j
.....···'11:;::;:1-0-
~~i:-:'l:::.Jx·
""'-:1= ••I"4T'
·r.... '· ....
tI
600 500 400
~
~
=
300
'J="~J';+: -j:,l
r,
rt~'·-I--t"""'
'it
L-
..
•
n
"llOUIO-
m.
"
80 - reNSlTY OF VAPOR - LeS/CU· FT. 8L fl L1QUIO" .. Uo-SUPERFICIAL VAPOR VELOCITY AT INITIAL FLOODING-FT/SEC.
+
S*-SURFACE AREA OF PACKING- SO. FT./CU.FT. TOWER VOLUME F*-F'RACTION OF FREE VOLUME IN PACKING
.002,
OF LIQUID - CENTIPOISES -GRAVITY CONSTANT-32.2 FT./SEC~
...
SHERWOOu
_......
~MlrLEY
..mI·;,·;:; I ;'i-;t:;:-..J:, AND HOLLOWAY. IND . ENG. DiEM. 30. 765 (1930) ~ rF!¥:tJ~ lii r:'~ •., !
·/+l!;I"·'
."
001
.01
.OZ
.03.04
.06 .08 .fO
.2
.3
I'
I
.,
.4.5.6.7.8.910
"'.
'n
100 80 60
.M -VISCOSITY 9
" .200
f"'";
SUPERFICIAL MASS vaoCITY OF VAPOR-L8SISEC/SO.FT.
2
3
·
~
AP/H