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
-
-
-
-
-
><
d
S ~
o
z00
2,2,3-Trimethylbutane (Triptane) ....................
-
20,620
19,110
25~
-
25"
-
20,570 20,570
19,080 19,080
25
0.237
20,550
19,060
515'
27"
-
20,540
19.050
50
51
.22
21,640"
20,290Q
196.5
45.4
.2:~3
21,040Q
19,690Q
5.00 5.22 5.08 4.99
293' 316' 310' 292.5
39" 37" 37" 39.5
20,840" 20,780" 20,750" 20,720"
19,490Q 19,430Q
19,400" 19,370"
.647 .661 .654
5.38 5.50 5.44
385" 398' 396'
36" 35"
20,710Q 20 ,660Q
19,360" 19,310"
35~
-
~O,640Q
19,290Q
84.5
.655
5.45
387"
36"
-
20,610 Q
19,260"
-292.0 -207.0
92.0 80.6
. 6:~3 .667
5.27 5.55
363' 401"
37" 35"
20,660Q
20,570"
19,310" 19,220"
146.4 155.4 154.2 153.7 154.6
-218.0 -231.0 -207.0 -211.0 -171
77.2 73.9 75.7 75.4 76.0
.678 .689 .683 .684 .682
5.64 5.73 5.68 5.69 5.68
463' 473' 472' 4i2' 473'
34" 34" 34" 34" 34"
-
20,450 20,420 20,400 20,420 20,400
19,100 19,070 19,050 19,070 19,050
30.1
-213.0
.595
4.95
249
70
-
20,880 Q
19,930"
50.5 24.1
-
83.5 94.2
.658 .627
5.48 5.22
343' 308
-
71.5 71.8 76.0
.697 .696 .682
5.80 5.79 5.68
420C
78.9 104
-234.0 -184.0
81. 3 82.9
.665 .685
5.53 5.70
350C
93.3
-231.0
74.8
.686
5.71
13.0
72.1
0.695
5.78
480~
243.8 245.4
-165.1
70.1 65.6
.702 .718
5.84 5.98
549< 551 c
114.2
228.4
-130
71.2
.698
5.81
530
C S H 18
114.2
210.6 -161. 2
71.8
.69u
5.79
OLEFINS Ethylene ........••.........
C 2H.
28.0
-154.7
-272.5
2n
. ;J;j
2.91
Propylene .......•..........
C 3H s
42.1
-
-301.4
140
.522
4.35
Butene-I ............. Cis-Butene-2 ......•...... Trans-Butene-2 ............. Isobutene ........ , ......
C,Hs C.Hs C.H s C.H s
56.1 56.1 56.1 56.1
20.7 38.6 -218.0 33.6 -157.7 19.6 -220.5
104 94.2 100 104
.601 .627 .610 .600
Pentene-1 (Amylene). Cis-Pentene-2 ............ Trans-Pentene-2 ............
CbH lO C;H1o CbH lO
70.1 70.1 70.1
86.2 98.6 96.8
-216.4 -290.2 -211.0
87.2 82.6 84.9
2-Methylbutene-1 ........... 3-Methylbutene-1 (lsoamylene) .......... " ..... 2-Methylbutene-2 ...........
C;H 1o
70.1
88.0
-
C;H lO C;H lO
70.1 , 70.1
68.4 101.2
Hexene-1 ............. Cis-Hexene-2 ............. Trans-Hexene-2 ............. Cis-Hexene-3 ............. Trans-Hexene-3 ......•......
C 6H 12 C 6 H 12 C 6H 12 C SH l2 C aH 12
84.2 84.2 84.2 84.2 84.2
C 3H.
40.1
-
Butadiene-l ,2 .............. Butad.iene-1,3 ..............
C.H s C.H s
54.1 54.1
+
Pentadiene-1.2 ..... : .. Cis-Pentadiene- L3 ........ Trans-Pentadiene-l,3 ........
C;H s C;H s C;Hs
68.1 68.1 68.1
112.8 111.6 108.1
-
Pentadiene-I,4 ....... 3-Methylbutadiene-l,2 ....... 2-Methylbutadiene-l,3 (Isoprene) ...................
CsH s CsH s
68.1 68.1
C6 H S
68.1
2-Methylheptane (lsooctane) . 3-Ethylhexane .............. 2,5-Dimethylhexane (Diisobutyl) ................. 2,2,4-Trimethylpentane ("lsooctane") .................
DIOLEFINS Propadiene .....•...........
C 7 H 1S
100.2
177.6 -
C SH 18 C sH 18
114.2 114.2
CSHl~
53.9
Heat of combustion as a gas-otherwise as a liquid. • Estimated. Q
-
-1.64.0 85.0
-
106
45
-
-
.234
-
-
-
-
410" 395'
420" 415'
Critical temperature-boiling point correlation. " Vapor pressure curve or correlation.
c
29.5
-
-
-
(1
o Z
lfJ ~
;>
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.
-<
.... (1 lfJ
-
20,230" 20,150"
-
~
=:r:
(J,j
trans-isomers.
PHYSICAL CONSTANTS OF HYDROCARBONS (Cont.)
FORMULA
BOILING MELTING MOLEC. POINT POINT WT. OF OF
DIOLEFINS (Cont.) Hexadiene-1,2 .... : ..•.•.... Hexadiene-1,3* ..•.••....... Hexadiene-l,4* ......•...•..
CSH lO CsH lO CaRlO
82.1 82.1 82.1
172 163 149
Hexadiene-l.5 ...•.•••...... Hexadiene-2,3 ..... : .•...... Hexadiene-2,4* .............
CSH lO CSH lO CsH lo
82.1 82.1 82.1
139.3 -221.4 154.4 176 -
3-Methylpentadiene-1,2 ...... 4-Methylpentadiene-l,2 ...... 2-Methylpentadie:le-1,3* ..... 3-Methylpentadiene-l.3* .....
CSH lO CSH lO CSH lO CSH 10
82.1 82.1 82.1 82.1
158 158.0 169 171
4-Methylpentadiene-1,3 ...... 2-Methylpentadiene-1.4 ...... 2-Methylpentadiene-2,3 ...... 2,3-Diroethylbutadiene-1,3 ... 2-Ethylbutadiene-1,3 ........
CsH lo CSH IO CSH lO CaRlO
82.1 82.1 82.1 82.1 82.1
169.3 - 94.0 133 162.0 155.7 -105 167 -
ACETYLENES Acetylene ..................
CzH z
26.0 -119**
Methylacetylene ............
CaH.
40.1
Butyne-1 (Ethylacetylene) ... Butyne-2 (Dimethylacetylene)
C.H s C.H s
54.1 54.1
~HIO
-
-
-
-
DENSITY °API
Sp Gr 60 0 /60 0
CRITICAL CONSTANTS
Lb Igal
OF
P Atm
-
t
-
64.5 67.8 70.6
0.722 .710 .700
6.01 5.91 5.83
495" 485" 470"
71.8 75.1 63.7
.696 .685 .725
5.79 5.70 6.03
454 475' 500"
65.0 67.0 63.9 59.7
.720 .713 .724 .740
5.99 5.93 6.03 6.16
4W'
480" 490" 495'
-
63.9 70.9 66.1 62.1 61.0
.724 .699 .716 .731 .735
6.03 5.82 5.96 6.08 6.12
490" 445' 485' 475' 490'
-
.416
3.46
94.9
.625
86.2 71.2
-
-
-
-
Gross
Net
-
--
-
103.5
62.0
0.231
21,47(}l1
5.20
275'
65~
-
20.810"
19.8W
.650 .698
5.41 5.81
375 420
65~
-
20.65QQ 20.51QQ
19,6oog
71.8 66.1
.696 .716
5.79 5.96
429 460"
-
-
g
-
20,45QQ
19,440g 19,340g
-
-
20.5W
19,390"
-
-
~
-
-
20,130
18,980
--
-
-
-
-
-
19,880
18.730
-
-
tj
>
~ 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/}
<
-60
-eo
3
2
4
5
7
6
8
?
cr
(.f-
MOLAL AVERAGE +-20
*
THE SLOPE AND AV. B.P. SHOULD BE DETERMINED FROM THE FOLLOWING FORMULAS: t,,70 - tlO S=
0
60
.,
-20
tv =
-40
4
IF THERE ARE INSUFFICIENT DATA THE 50% POINT MAY BE USED FOR THE VOL. AV. B.P.
-60
FOR WHOLE CRUDES:
-80
tv: -100
-120
-140
tlo+2t50 +t 9O
2
3
4
5
6
7
15
8
'.
t30+t.50+t70
3
(
;
CHARACTERIZATION FACTOR VS
14.0
BOILING POINT AND GRAVITY
13.0
13.0
12.0
12.0
_ 11.0
11.0
~
10.0
10.0
90
9.0
100
200
300
400
500
600
800
700
..
900
1000
I 100
..
200
CRUDE
13.6
•
13.2EEHHl 1 m:aE 3.0
12.1:)
a:m:a
400 .
TYPICAL GROUP WHITE PRODUCTS
134HHH1
300
a
I I II II
I-II
II II-III
IRAO IRANIAN EAST TEXAS S.LOUISIANA .lUSEPIN
II II
II-III II- III
III
III
WEST TEXAS TIA .lUANA (NED. COLOMBIAN LAGUNILLAS
liI
III
v
V
III III
a
102)
III IV
600
700
~
900
1000
;.t~ 1 I:n __
CHARACTERIZATION FACTOR ~ BOILING
I
I I
. .-!{ ~~ili ~ r .1%1+I.l-++J.1=:+I::t1+
POINT
TVPICAL CRUDE
~m
FRACTIONS
13.6 13.4
13.2
T·
II II
>-
r
IV IV
-...
13.0
•
12.6
-
800
.,
GAS OILS HEAVIER
PENNSYLVANIA RODESSA PANHANDLf MID' CONTINENT KUWAIT
500
,
.
:i '.
12.8 12.6
>-
I' .
,
12.4
12.4
, .of:
12.2
12.2 .~
12.0
12.0
11.8
11.8 r
11.6
II ~
tt
r
11.4
TIT
"
d~
..,.+1.t
/00
200
300
400
500
600
I
., .
r
800
. i. _
' 1->-
I·~ -I.
700
1/.4
,
i
r i
11.2
H
"
r.
11.6
,J;
_Ir H-I" ;'++1+1-++1 11.2 .,:
900
1000 I
200
100
300
400
500
600
700
800
1300
JOOO
GRAVITY ~ BOILING POINT TYPICAL CRUDE FRACTIONS
70
60
60
50 IHB111lHlIIII1l1l1J 1111 II tlHIIJIlJI IlIllIIlIIlIHt"KINI!'N:"lIIJilJIIH1I1lIlit11J»111 1111111111111111 111I111l"KlOIIJ II i1II1III1II1IJIIJLLIUHlIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIII11II11111111 flll150
....
00
40
40 CRUDE
TYPICAL GROUP WHITE PRODUCTS
30 I:fI::i::l:!:1I
20~
10
PENNSYLVANIA RODESSA PANHANDLE MID· CONTINENT KUWAIT
WEST TEXAS TIA JUANA (MED. COLOMBIAN LAGUNILLAS
10J
I I
II II
% II tlllllllllI11tllltul1tlJIlItn1'W"~1i\IM1t1n&J~IIIIIIIIIIIIIIIIIIITI't'l+l>lilJ II II· III
II
a
102)
200
GAS OILS HEAVIER
I I.
I·II
IRAO IRANIAN EAST TEXAS S. LOUISIANA .IUSEPIN
a
FEIBHtlilMIllllllllllll130
II·III
II III III III
II-U.I II
III III. IV V
nI IV
II
In
II1II111111 i 1IIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIIImmmiiIII III IlrmUWOU1l11I1't'ffi!lratJUI:I!li1ln'iJmBTBI20
I1I'
V
300
400
500
600
700 ..
800
900
1000
Section 3
MOLECULAR WEIGHT The molecular weight chart for pet.roleum fractions on page 21 was derived from an empirical correlation of molecular weight and the function, T,,Jso.4o,· where T m i" the mean avcrage boiling point of the fraction in oR, and s, the specific gravity at 60°F/60°F. The ayerage deviation for about one hundred petroleum fractions from 75 to 500 molecular weight is +20/0. Up to a molecular weight. of about 300 this correlation applies equally well to pure hydrocarbons, with the exception of normal paraffins, which have lower molecular \\'eights than predicted by the chart. Above 300 molecular weight most pure hydrocarbons for which data are available deviate from the correlation in the same direction as the normal paraffins. An explanation of this incongruity may be that these particular high molecular weight compounds have relatively long chains and consequently should fall somewhere between the normal paraffins and the multibranched and multicyclic hydrocarbons in petroleum fractions. The molecular weight of crude fractions is given as an independent function of mean average boiling point, page 22, and also of gravity, page 23, for approximate use when only one of these variables is known. Examination of these charts shows that the boiling point chart is much less susceptible to variations with type of crude than the gravity chart and, consequently, will usually give a better approximation than the latter. However, in general, gravity rather than the boiling point will be available.
GENERAL REFERENCES API Research Project 42. Bridgeman, Proc. API 10, No.2, p. 124 (1929). Doss, "Physical Properties of the Principal Hydrocarbons," 4th Edition, The Texas Co., New York, N.Y. (1943). Fitz imons and Thiele, Ind. Eng. Chem. (Anal. Ed.) 7, 11 (1935). Francis and Wood, J. Chem. Soc. 48, 1420 (1926). Kay, Ind. Eng. Chem. 28, 1014 (1936). Mail' and Schicktanz, J. Research Nat. Bur. Standards 17, 909 (1936). Mail' and Willingham, .T. Research Nat. Bur. Standards 21, 535, 565, 581 (1938). Rosenbaum, J. Chern. Phys. 9, 295 (1941). Shepard, J. Am. Chern. Soc. 53, 1948 (1931).
19
~
400
600
500
600
700
800
900
1000
1100
MOLECULAR WEIGHT n BOILING POINT
600
NORMAL PARAFFINS AND ISOPARAFFINS
.
500
500
400
400
200
200 180 160 I - AVERAGE OF ISOMERS CONTAINING A SINGLE METHYL OR ETHYL BRANCH 2 - AVERAGE OF ALL OTHER ISOMERS
140
120
120
100
100
80
80
60
60
40
40
20
20
-100
a
100
20
200
300
400
700
800
1000
900
1100
1200
MOLECULAR WEIGHT Yo! BOILING POINT AND GRAVITY 600
PETROLEUM' FRACTIONS
500
400
300
300
280 260
200
240 220
200
200
180
180
160
160
140
140
120
120
100
100
80
eo
100
200
300
400
500
21
600
700
aoo
200
400
300
500
700
600·
800
900
460
460
. MOLECULAR WEIGHT ~ BOILING POI NT 440
TYPICAL CRUDE F RACTIONS
440
420
420
400
400 CRUDE
TYPICAL GROUP WHITE 'RODUCTS
380 360 340 320 300
PENNSYLVANIA RODESSA PANHANDLE MID - CONTINENT KUWAIT
a
GAS 01LS HEAV lEA
I I
I I. II
360
II
II-III I I -III II-III
III
II II
UI
In
IY Y
IV
I-II
IRAQ IRANIAN UST TEXAS S. LOUISIANA JUSEPIN
II
WEST TEXAS TIA JUANA (NED. COLOMBIAN LAGliNILLAS
nI
340
In
III 102)
.,
I
II II
a
380
III
320
IV
300
Y
280
280
260
260
240
240
220
220 CHARACTERIZATION
VISCOSITY
INDEX
200
GROUP
1
12.1 -12.6
SO-IOO
,eo
II
11.9 -12.2
60-S0
III
11.7-12.0
40-60
IV
11.5 - II.S
20-40
y
11.3-11.6
0-20
160
FACTOR
* DEWAXEO
OF LUBE FRACTIONS·
200 180 160
TO +20 o F POUR
140
140
120
120
100
100
200
300
400
600
500
22
700
800
900
10
20
30
40
50
70
60
MOLECULAR WEIGHT
460
~
BO
GRAVITY
460
TYPICAL CRUDE FRACTIONS
440
440 CRUDE
TYPICAL GROUP 'WHITE PRODUCTS
420
360
. WEST TEXAS TIA .lUANA (NEO. COLOM81AN LAGUNILLAS
340 320
GROUP
IX
I-U II
11
III
IV
280
y
*
260
400
II-%II
380
lI-nI
II-XII II
II
Dr
III
102)
% :II
:III
II III
UI
XII
DI IV
IV
IV y
v
~
360 340
CHARACTERIZATION VISCOSITY INDEX FACTOR OF LU8E FRACTIONS
I
300
I
II
a
420
~
~
IRAO IRANIAN EAST TEXAS S. LOUISIANA .lUSEPIN
380
GAS OILS HEAVIER
I
PENNSYLVANIA RODESSA PANHANDLE MID' CONTINENT KUWAIT
400
a
320
*
80-100
12.1 -'2.6 11.9 -12.2
60-80
11.7 - 12.0 1/.5-11.8 1/.3-11.6
40-60 20-40 0-20
300 2BO
.
DEWAX£O TO +20 o F POUR
260
240
240
220
220
200
200
180
180
160
160
140
140
120
120
100
100
110
20
30
40
50
60
23
PIA
PABLO
70
80
o-;n'i L ,
If"
\)
_,_
f.J . .JJ
t. I.,.
1VI0T'l'A
Section 4
VAPOR PRESSURE In developing the vapor pressure curves for most of the individual hydrocarbons, the reciprocals of the absolute temperatures were plotted against those of a reference compound (ethane, butane, or hexane) at the same vapor pressures. 1 With one or two exceptions, this relation was linear over the entire range of the data, but if a slight curvature was indicated, as in the case of benzene vs. hexane, a straight line was not imposed upon the data. The vapor pressure curves for methane and the reference compounds were developed directly from the data by plotting vapor pressures against reciprocal temperatures. Most of the reliable data fell within -I- 1OF of the correlations, and this may be considered as about the accuracy of solid portions of the vapor pressure curves. Normal boiling points in all cases were taken from "Selected Values of Properties of Hydrocarbons."2 While vapor pressure is meaningless above the critical temperature, the curves were extrapolated beyond this point so that other properties in the liquid phase could be calculated in the absence of any other data. For example, these extrapolated curves may be used to make rough approximations of the fugacity, density, and enthalpy of hydrocarbon vapors in solutions at temperatures above the critical. The generalized vapor pressure charts for hydrocarbons were also derived from the linear reciprocal temperature relation with hexane used as the reference compound. The pressure scales correspond to the vapor pressure of hexane as a function of reciprocal temperature. The temperature scales were based on the reciprocal relation up to 700°F, but above 700°F it was necessary to modify the scale to secure better agreement with data on high boiling hydrocarbons and petroleum fractions. 3 The slopes of the normal boiling point lines on the rectilinear chart and the corresponding points on the alignment charts were based on normal paraffins. However, with the exception of some of the lowest boiling members of the various series, there is a good indication that these charts apply to hydrocarbons in general. In API Research Project 42, the boiling points of a large number of 1 This is the most nearly linear of the simple vapor pressure relations, with the exception of a similar one where the reciprocal temperatures are plott.ed at the same reduced vapor pressures. 2Nat. Bur. Standards Circular C461 (1947). 3 Beale and Docksey, J. lnst. Petro Tech. 21, 860 (1935).
24
VAPOR PRESSURE
25
different high boiling hydrocarbons were determined at 0.5 mm, 1.0 mm, and 760 mm, and these were checked against the low-pressure alignment chart. The average deviation was about 2°F over an average extrapolation of around 400°F, and there was no trend between the paraffins and other hydrocarbons. Thc cxtrapolation of the vapor pressure scale below the hexane data has been checked indirectly by the Clapeyron equation using thermal data on hexane at low tempcratures. Also, low-prcssure data (below 0.001 atm) on petroleum fractions are in good agrecment with this correlation. GENERAL REFERENCES Aston, Kennedy and Schumann, J. Am. Chem. Soc. 62, 2059 (1940). Aston and Messerly, J. Am. Chem. Soc. 62,1917 (1940). Beale, J. Inst. Petro Tech. 22, 311 (1937). Beattie, Hadlock and Poffenberger, J. Chem. Phys. 3, 93 (1935). Beattie, Poffenberger and Hadlock, J. Chem. Phys. 3, 96 (1935). Bea.ttie, Simard and Su, J. Am. Chem. Soc. 61, 24 (1939). Bea.ttie. Su and Simard, J. Am. Chem. Soc. 61, 924 (1939). Bekhedahl, Wood and Wojciechowski, J. Research Nat. Bur. Standards 17, 883 (1936). Benoliel, Thesis, Pelillsylvania State College (1941). Benson, Ind. Eng. Chern., Anal. Ed. 13, 502 (1941). Brown and Coa.ts, Univ. of Mich. Res. Circ. Series 2 (1928). Comrp.unication from The ::\1. W. Kellogg Co., New York, N.Y. Dana., Jenkins, Burdick and Timm, Refrig. Eng. 12, 387 (1926). Doss, "Physical Constants of the Principal Hydrocarbons," 4th Edition, The Texas Co., New York, N.Y. (1913). Ega.n and Kemp, J. Am. Chem. Soc. 59, 1264 (1937). Francis and Robbins, J. Am. Chem. Soc. 55, 4339 (1933). Frolich and Copson, Ind. Eng. Chern. 21, 111G (1929). Garner, Adams anu Stuchell, Refiner 21, 321 (1942). Hei ig, J. Am. Chern. Soc. 55, 230-:1: (1933). Heisig anu Davis, J. Am. Chem. Soc. 57, 339 (1935). Heisig and Hurd, J. Am. Chem. Soc. 55,3485 (1933). Ingersoll, Thesis, Mass. Inst. Tech. (1930). Intel'l1ationa.l Critical T:1bles, Vol. III. Kassel, J. Am. Chem. Soc. 58, 670 (193G). Kay, Ind. Eng. Chem. 30, 459 (1938). Kisti:1kowsky and Ricc, J. Chem. Phys. 8, 610 (1940). Kistiakowsky, Ruhoff, Smith and Vaughan, J. Am. Chem. Soc. 57,876 (1935); 58,146 (1936). Krase and Goodman, Ind. Eng. Chem. 22, 13 (1930). Lamb and Roper, J. Am. Chern. Soc. 62, 806 (1940). Kinuer, J. Phys. Chem. 35, 531 (1931). Livingston and Heisig, J. Am. Chem. Soc. 52,2409 (1930). Loomis and Walters, J. Am. Chem. Soc. 48, 2051 (1926). Maxwell, Ind. Eng. Chem. 24, 502 (1932). Morehouse and Maass, Can. J. Research 5, 307 (1931); .11, G37 (1934).
26
DATA BOOK ON HYDROCARBO S
Nieuwland, Calcott, Downing and Carter, J. Am. Chem. Soc. 53,4197 (1931). Pitzer and Scott, J. Am. Chem. Soc. 65, 803 (1943). Rintelen, Saylor and Gross, J. Am. Chem. Soc. 59, 1129 (1937). Sage, Lacey and Schaafsma, Ind. Eng. Chem. 26, 214, 1218 (1934). Sage, Webster and Lacey, Ind. Eng. Chem. 29, 658 (1937). Schmidt, Thesis, Paris (1934:). Stuckey ll,nd Saylor, J. Am. Chem. So:;. G2, 2J~ (1940). Vaughan, J. Am. Chem. Soc. M, 3863 (1£::>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-=~. .
_--<
200
=:h:b-~.1-'-.
~=t=t::= ~=;::
H
VAPOR PRESSURE OF HEPTANE
.4
.3 ~t:
.2
200
~~"'~~'t:g~E~ 'fg -= .
.:
~."
70
60
=t...:
50
40 30
20
H-H-++-H-t--t-++-t-f-fH--f-+-HH-+++! =f:'~= ~", ~::.~f ~;; -:~-=:.=
'C, E.: ~! :t+' - §j:, = ~
- -. '._. ... .-
10 9 ,..:::::: 8
:;:::=.;.
.008
.em .006'
.005
f-:-. :-'~"'-""::l:
.i=i= I/.·~ . .::t± t':i::~.
l::::tl"
.0
.003
~-
..:p.
..... :
I-:=±: '-.q:: --l-
'-;- 1-.; ':-i-:'-
";J-
. c;::::= c-F-¥+'·~f-C.,....=· r_ L:.I ..
.= ..
7
.1=;""<- H
::-J: :.
6
tH--
5 4
. ·+--'---~"'H
Xt__ 1--
_.-.
.. !-' . ; .
3
. '--.-1-1- ..
_J
.002
33
2
50
1.0 .9
200
ISO
100
250
=I-·,~-,,::'~,':7~,£::I-,-:.-~:: -~~
.8 .7
=F~:::l~+:-
-
..: .
-f-
.6
.5
VAPOR
:4
OF
PRESSURE OCTANE ~'.
-I- '
--;
':,::J
.2
,'+'++++++++++++++++H-J-H " ...:' - t -
•
_
.01
100 90
70 60
.06
.05 .04
:
"
'I
.0/
A
.009
, ,
-1---J---~--1-I-I-
.009 .007 DOG
.005
I-
J-+-tft-t-H-t-t f-
"
i
I
200
i
300
500
400
34
600
700'
800
:1..-=
VAPOR PRESSURE OF C3 UNSATURATES 3
.
,~-.
-'-"
-,
'-~:j=~~
-
-Ej
-:1# ., -
~I= --,..~ r'
.1-/
_I-~-~.
to
300
t-:~-
j-++
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I
200
WW-
1- 1
.J
I
1
.1'".
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I
l'i
o
I
I
I
100
l'
I
300
200
35
41)0
500
()
-!'>("}
_'00
1.0 .9 .8
-
50
.7 .6 .5
VAPOR PRESSURE
.::
OF C4 UNSATURATES
.3 , -I
200
.2
~
I
''1
I
.1 .09
.-
-
- =Z:~-E§,.;:.
-:'E
~, ~----.
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t'-
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I
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.
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COMPOUND
11
CRITICAL PT. of ATM.
1,3 BUTADIENE
24.1
308
45 .
VINYLACETYLENE
42
365*
75 v
ETHYL ACETYLENE
47.7
DIMETHYL ACETYLENE
80.4
375 420
65 v 60 v
* ESTIMATED
v
VAPOR PRESSURE CURVE
I 9
8 7
6 5 4
3
2
I
I
o
II
100
300
200
36
400
500
-.
1.0 ,9 .8
.7
100 .
=r_0-
150 .
..
•
.5 .4
.
0
---
.
VAPOR PRESSURE OF BENZENE AND TOLUENE
.6
..
200
.... .
0_
.
o~
~=
--
-
.~
.2
.t
.-
..
.09
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I I
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.
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ti
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:
I I
oir:'1
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-
. . 300
I
I
400
37
500
600
700
50
1.0 .9
-.- . - .
.8 .7 .6
100
.
-
..
.
..
.. _. .
-~-
M
-
-
.4
B.P.
CRITICAL POINT
~
~
ETHYL8ENZENE p-XYLENE m-XYLENE o -XYLENE
.2
250
..
-
271.1 281.0
655
38
652
35 V
W~ i=1-1=l=1=!.
r
~
..
:+ .
--
=.l
..AItL
300
'-
-
=).
r-'
COMPOUND
.3
2o0 -
'.
VAPOR PRESSURE OF Ce AROMATICS
..
.5
150
--
-
--J
282.4 655" 36 V 292.0 675 37
CURVE *ESnMAlD1 V- VAPOR PRESSURE . ,
.1 .09
I
I
I
I
I
I I I
..--.. ,
.. "
I I I I
I I
I I
I I
I I
I I
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L
,
I
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..
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o 5o 4o
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30 ..
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.
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VAPOR PRESSURE OF P - XYLENE EQUALS VAPOR PRESSURE OF ETHYLBENZENE MULTIPLIED BY .950
,
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600
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VAPOR PRESSURE
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OF CYCLOPARAFFlNS ~
-
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39
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VAPOR PRESSURE OF HYDROCARBONS
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01
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VAPOR PRESSURE OF HYDROCARBONS
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v 50
100
150
200
250
300
350
A
400
'/
450
500
tt1
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550 ----600 650 700
i I I "
800
I
900 1000 1100 1200
NORMAL BOILING POINT .
t. ,
,I
Y! BOILING POINT AT 10 MM
."i,
.1.-
"!
1100
1000
900
800
700
600
r
500
400
100
200
300
400
·500
600
43
oPI
700
r)H,T 1r~T _ TJ) ~ i tV/l. \. J .)
PABLO MOTTA
900
VAPOR
PRESSURE OF GASOLINES
180
.7 .8 .9 1.0
170
2.0
70 80 90 100
3.0
150
4.0
200 ci J:
* CURVE SLOPE OF DISTILLATION + LOSS (A.S.T.M.). ° F@ 15'>'. - of 5°4
150
1.5
~
10 IN THE ABSENCE OF DISTILLATION DATA THE FOLLOWING AVERAGE SLOPES MAY 8E USED:
140
LIGHT NAPHTHAS (F.B.P.-300°F) NAPHTHAS (F.B.P.-400°F) AVIATION GASOLINES MOTOR GASOLINES
130
°F/'>'. 2.5 4 2
01234 ~
3
2
0 2 en "(Ij 3 Q)
110
3
I
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:::>
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0: L&J
f
12 14 Q:" 0 16 Q. § 18 20
::E
...
l&.I
70
0
W
a:
60
l&J
6 7 8 9 10
UJ II
0..
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5
....
u.
0:
I
6.0
0:
w
0..
8.0
:::>
...
.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
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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).
~
..
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FUGACITY FUNCTION
1m
OF METHANE
_. t
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200
r v.
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10
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I
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100
90
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80 70
60
50 40 30
20
-
I
: :.-;;!
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_
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•
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-
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10 9
8 7
6
5 ,
4
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3
-300
-200
o
-100
49
100
200
300
&I .
..
F~~~CITY FUN~TI~N OF ETHYLENE
I
r +-
-
..
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1 00
90
80 70
60 50 40 30
20
-
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_I~ 8 7 6
5 4
-200
o
-100 50
100
200
300
~
-~=
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-
--- --
-
-
---
=
= = = = = -
FUGACITY FUNCTION OF ETHANE --
-
~
00
•
--
--
-
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-
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-
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-
-
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I 00
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-
90
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-
80
-
70 60
--
50
40 30
20
~
-
-
~
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10
-'
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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
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-
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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
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0.4
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0.3
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i
90 80 70 60 50
40
-
30
20
11 -(
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100
53
I I
200
300
400
__
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~
-
r
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90
80 70 60
,-
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30
20
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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'
H~t-H
!/
II ,"
I -
::E-
,-:=
o-
_~c:~~'%
.)e;d _.
:-r' '. ';:
~ ~~
~~.
.~' ~--
/~ :-='/--::- r:t=-1='
E' r-+-'
'~.:i=d-. ~ ,.-.
.,;-
:~~ 10 9 8 7 6
5
, 1=
,
4 H-
-:!=.::l~·::r!.::
.e.;-::!::C
-+- ~~.~!-:. -. --
:::c.J'
3
1""- -
1-1-
2
a
100
200
57
300
400
500
o
100 ....
0.9 0.8
200 .
::E;'?:::f ~i=ft.
± .~ E
-I/~~
I....:_!c·~
E'-=H'
0.7 0.6
(
.
.
.CE1.
FF'
.::t::;::
,::
:t
~
0.5
FUGACITY FUNCTION OF HEXANE
~/~ ...
0.4
= ......
:~ :4'; i'c
• ::::
= ....
t-
0.3 ,--'
0.2
I I
c..,
..
.£ -:- ...... -
•
~IOO ~90
'. .
~
-
80 70 60
~-."'_
~t--
"-
..=l=E ~--::-"'::: . . . . ~'-, _., ....
-=..-_"-
--4-
'-!-
''-
"""T
50
-!--,I o
40
-"'-r--
._
30
20 Ii
I
I I
.I -
:t:-.t=:
_ :-~ .;:~ •. ::..:
.:
;:::;:-~. ::H:~;:-: . =r-r :-E~.·U·~· -r:±: . -0-'1:';=1....".[ .:~t""e§::l=: :t=t:: ...c.. r=::::t:::::t=t:
=r-
.I~
'-i-:;:-_
"
e
~~
=j---:-'
7 6 5 4
3 ,
t:b::
2
...
! I I
1/-
, I
0
-
V
I
100
200
58
300
400
500
200
300
O.9~
0.8 0.7
;::.
0.6
§ ~ ~
0.5
I=l=I
FUGACITY
FUNCTION
OF HEPTANE
= -
0.4 ~.
-->-,-
0.3
~
-
-'f
-8
-
0.2 ClhHH-+-++++-++-H-Il--JI+-H--H--I-HH-+-++-
'T
-
II
I
0.09 0.08
'
0.07 0.06
o
0.05
o
0.04
o
100
I.O~
0.03
20
0.02
4
3
2
~
!/
..
1I
200
100
59
300
400
500
100
200
300
0.8 0.7 0.6
0.5 0.4
~
0.3
FUGACITY FUNCTION
E
0.2
OF OCTANE
1~ :
I
I
0.08 0.07 0.06 0.05
0.04 0.03
30
0.02
20 ++++H-H··t-H-+-++ I-f--H-H-+'++++++-++I-I-f--H-+--J----I--++++++-hj£.j-I-I-hjoq~ -+--I--t-+-t-+++-+-+-+4 '
-t'i-"'Fff J.
-1:. t=
.-'-= I.
+ie,'}l
t=:± =Lfi f-L - F.'r
.f-
::-~
1
&
...
I-~
~ - f:::j:. - -' ,-, . • +-t-
4
- ~~rj' ~ ±-l-I' -
3
2
II
100
200
GO
II
300
400
500
FUGACITY FUNCTION OF HYDROGEN
p ::;
THE FUGACITY FUNCTION OF HYDROGEN. IT"fplfll'' IS BASED ON A PARAFFIN SOLVENT HAVING A MOLECULAR WEIG HT OF 114 (OCTANE). ~ FOR OTHER SOLVENTS MULTIPLY THIS FUGACITY FUNCTION BY THE ..... CORRECnON FACTORS, A, FOR MOLECULAR WEIGHT, AND FOR CHARACTERIZATION FACTOR OF THE SOLVENT~
P
e.
tHw
THIS CHART DOES NOT APPLY AT TEMPERATURES GREATER THAN 0.95 TIMES THE PSEUDO-CRITICAL TEMPERATURE* (OR) OF THE LIQUID PHASE. T pc: XH (60) + X HC (T HC )
*
I
I
I
I I I
o .
--
I
_.. - __-··_··-0
. I
-r
,
==._.
~.
2000
--
,
~
K:l
50
100
150
200
250
300
3 1.0 0.9
0.8 0.7 0.6 40
61
....
60
80
100
120
I
140
IGO
.2
.3
.4
.5
.6
.8
.1
9
10
11
12
13
1 ~.O
.9
9
.8
.8
.7
.6
.6
..
.5
T-
= A fLr
II-
f p IP
"':~-:.
-
=-
~~
.5
.
VS PIPe FOR LIQUID PHASE
flT/IT VS IT/Pc FOR VAPOR PHASE
.S i'
FUGACITY OF HYDROCARBON VAPORS
lJ
.2
h
-,
,6.. -;fu. ~ tl:f .1·...,;....+..........."
It
~
.1
+.-!
LEWIS AND KAY. OIL AND GAS J. 32. NO. 45, 40 (MARCH 29, 1934)
RE~ERENCE:
o •
.1
.2
.~
.5
.6
.7
62
.8
.9
1.0
J.2
1.3
1.4
',4
4 I. 0
1.6
1.8
2.0
2.2
l.4
2.6
2.8
3.0
32
3.4
3.6
3.8
FUGACITY OF HYDROCARBON VAPORS .9
.
~
.. .9
.8
.7
.7
.7
.6
.6
.6
.5
.5
A
.4
.2
J
ot.4
1.6
1.8
2.0
2.2
2.4
2.6
2.8
63
3.0
3.2
3.4
16
18
4.0
o
<-
RELATIVE VOLATILITY OF LIGHT HYDROCARBONS
2.00
1.80
1.70 .-++-
.~
1.60 ETHYLENE (ETHANE
1.50
1.40
-::1:; .
,
-_l
::l"-
~+
-
1.30
,
1.20 1.10
:r
1.00
-100
100
0
200
300
400
:h
1.40 1,30
PROPYLENE/PROPANE Ir.
1.20
1
,
~
-
.
1,10 1.00
•
-100
0
100
200
4C0
RELATIVE VOLATILITY ;OF LIGHT HYDROCARBONS
1.30
I J
ISOSUTENE/SUTANE FOR BUTENE °l/BUTANE MULTIPLY BY 0,980
1,20 ct' I-
1I0
1.00
o
100
200
300
400
TRANS-SUTENE-2/SUTANE
.90
1+=
o
1...
J..~
r
100
200
300
400
CIS - BUT ENE- 2/SUTANE
1.00
• .90
.eo
0
100
200
300
400
1.3-SUTADIENE/SUTANE
tOO LIO
tOO .90
0
100
200
300
65
400
•• 1
.
."
, :'-' -!:-i1-
RELATIVE VOLATILITY +
,.
80
1111
I.IO.
.
.
o 66
~ ~O
.
1111 400
8
9
~~
~~'I:
= FUGACITY CORRECTION FACTOR LIGHT HYDROCARBONS IN ABSORBER OILS
_ t- .
2.5
2.0 - ~ L :9..
MULTIPLY FUGACITY FUNCTIONS (OR VAPOR PRESSURES) OF LIGHT HYDROCARBONS BY CORRECTION FACTOR WHICH IS INDEPENDENT OF TEMPERATURE AND PRESSURE
1.5
*CHARACTERIZATION FACTOR OF THE LIQUID PHASE IS A WEIGHT FRACTION AVERAGE OF THE CHARACTERIZATION FACTORS OF THE ABSORBER OIL a DISSOLVED HYDROCARBONS.
REFERENCE: COMMUNICATION
FROM lHE M.W. KELLOGG CO .. NEW YORK. N.Y.
,-..
8
9
10
"
12
67
13
Section 6
CRITICAL PROPERTIES Analogous to PUl'C substanccs, thc true critical point of a milltUl'c is a uni'lUC point on thc phase cnvelope where thc dcnsitJy and composition of the vapor phase arc identical with those of the liquid phase. Sincc the compositions of the two phascs arc the samc, fractionation of a mixturc is impossible at the critical point. Conscqucntly, the dcgrec of approach to thc critical point of a mixtuTe sometimes serves as a rough guide to thc fcasibility of separating thc components by fractionation. For PUl'C hydrocarbons, it has been found that a number of physical propcrties may bc correlated by reduced tcmperature, TIT e, and reduccd prcssure, PIPe. Various data have shown conclusively that none of these correlations apply to mixtures if the truc critical temperature and pressure of the mixture are uscd to determine the "edueed conditions. This difficulty has bccn overcomc by thc introduction by Kayl of thc concept of pseudo-critical tcmpcrature and pressure. By using thc pseudo-critical tempcrature and pressure to predict the reduccd conditiuns, Kay found that compressibility data on pure hydrocarbons could be applied to mixtures. Although Kay determined the pseudo-critical point by averaging the critical properties directly for known mixtures and from the averagc molecular weight for pctrolcum fractions, it has been found that much bcttcr results can be obtained by using the average boiling point method proposed by Smith and Watson. 2 As Smith and Watson pointed out, the true and pseudo-critical points must approach each other as thc boiling rangc of a fraction approaches zero and must coincidc for purc compounds. These conditions arc fulfilled by the charts in this section applying to petroleum fractions. Smith and Watson's relation between t11C true and pseudo-critical pressurcs on page 74 has been checkcd by the true critical data of Kay on ethane-hcptanc 3 and cthane-butane· systems. These data confirm Smith and Watson's eurvc well into thc region of their recommended extrapolation.
GENERAL REFERENCES Doss, "Phyaical Properties of the Principal Hydrocarbons," 4th Edition, The Texas Co., Nel\' York, N.Y. (1943). Internationnl Critical Tables, Vol. III. Roess, J. Insl. Pelr. Tech. 22, 665 (1936). 'l{ay, Ind. Eng. Chem. 20, 1014 (1936). 2 Smith and Watson, Ind. Enu. Chern. 29,1408 (1937) . • Kay, hId. Eng. Chern. 3D, 459 (1938) . • Kay, Ind. Eng. Chern. 32, 353 (1940). 68
,
800
._ m Il ~JI '
900
1000
m
1200
'I
.
°illii
"
.
600
..
I i
,
500
ti, rYl·j I'~! -
0
In 1m
r:. iI!lt-f :-!, .-. T' r ", ljl' ,
,
100
f!
"
"
,
, I " :r
I
'J
I
r; I ' J , I, I
11
,,
.
;.
,
j ,
j'
.
I
,j i •
I
,
I
rnu:
.
Ill~
,
mm
,
,,' I
I it
I.11
I
"
m
, ,
I 1
'i
!WJ
· I
·
I"
.1.: . I j'" 1-1' I . . ' t
1t .
It
,I
1
,
.,
I
I'
!II
.
!l
f
f
• J
L
I !
i
I
!
i
,
I
f J!
IIIl
..
r
I,
·
,
I
, I
j
200
I
I
.
1
,
,
Iii
'lil!1
I ' Hli' ,: ! HI' , I 300
.
I
·
I
,
,j"
'II!, IllliQ,HHlii!!:
400
500
69
600
t
.1
I II i" ::' J! • r I ! • II I II j f, I t ,! i' ill dlj!1! liid!! tl! : .. :-1"' . I, ~.1. ;~: I r,I! tl HI ':: IJ I 'I. ,I', i'l:1 II i' !I 'l f ~ ,.. ~:; ." rr"l , ill ,'f· ''I '1 ' ~.,.., :-1'; ;, ! , 1. ' .. I " , . . , ! ' 1.I 1111["1 It!; I fH. t " . i jiPi ' :. f 1 ,,' ! ' r_1 :. •L 1! ,I '. I ,! t:tll. . t: fl lill Ijl ill lit!, . iu,1• liP 1f : t . Iil" !. ,'I"t. :1:: 11 ,•. 1tI I , IIU11
inl
Iim1
mrr
1
"
f
,.
I I
,
t
·
r, (1 ,
tlftfi, , , .
,
~
.
i- t1 '.
r
•
IIWJ
I
!
.
r• ~; 111'
,
mitIHI I
,
I~
Hm1
11m!
t
J
70
j~
lOOl iMj
,
.
· ~
l1m1 ·
I
800
•
m
•
,
900
I
11
I
1000
II
CRITICA L TEMPERATURE OF PURE HYDROCARBONS
I
m
1100
I~
ffim
I!PI : I
,
III li1
700
•
Jfi1,
,
.'II 1'0 1 tub
800
-220 -200 -190
80
-160 -140
., 60
440
CRITICAL TEMPERATURE 40
OF LIGHT HYDROCARBONS
q,
PURE
420
COMPOUNDS AND MIXTURES
20 -
-.
-
.f!:[ ,-
400
'l* ... ' .-
+
t"
i ,
u
0
390
,
...L
-20
360
-40
340
-60
320
-80
300
·100
290
-120
260
r
240
j.
220 [
, .. I
1r
I ..
In I
"
I"
I i:
1
_I
t
I J I
I
200
THE BASE CURVE REPRESENTS CRiTICAL TEMPERATURE VS. BOILING POINT fOR PURE HYOROCARBONS ANO PSEUOO-CRIT'CAL TEhIPERATURE vs. MOLAL AVERAGE BOILING POINT FOR MIXTURES. THE GRAVITY CURVES REPRESENT TRIJE
I ,
I
CRITICAL TEMPERATURE YS. WEIGHT AVERAGE
190 160
BOILING POINT FOR MIXTURES. fOR ALL HYDRO-
CARBONS THE PARAFFIN GRAVITY OF THE SAWE BOILING POINT 5H:XJLO BE USED IN
I I
140
COMPUTING THE GRAVITY OF THE MIXTURE.
., I .,
:_i¥ill -140
l
I
t;
~
If
~
I
I
I~
I
[
f
,I·
120
.fiT • 1h . ..
I
100
, I
j -,
80
1
-120 -100
-80
-60
-40
-20
0
70
20
40
60
80
100
120
140
/
!mI
~
~
-
CRITICAL PRESSURE OF NORMAL PARAFFINS
20
'~
60
40
80
100
120 i50
Ii 40
40
30
THE PSEUDO-CRITICAL PRESSURE Of LIGHT HYDROCARBON MIXTURES HAVING AN AVERAGE MOLECULAR WT. LESS THAN BO CAN BE DETERMINED FROM THIS CURVE.
--.-:
.
20
.-:-
..
20 I
I
.
-
15
-
J
I
10
140
t+:
-
Iii L
-
.
-
r-
i~- ~i 160
.
180
200
220
71
240
260
280
900
§
800
~
1300
t ....
" .....,....,
1000
..=r::
':Ct 8
-1
.... ,
",.
>:.
~
.
~".
'- "-J :.... ,
..
ct= •
.. -
...
1100 ;.~r:$·
...
:$. • t
,-
"
~
=
...
800
'1'=: •
,-
700 .
.::r; ":-1.ir
800
,-
.... ..
--
~
.
.~
100
200
300
400
500
72
600
700
800
100
200
400
WO
500
100
7('.
r
-
PSEUDO- CRITI CAL PRESSURE OF PETROLEUM FRACTIONS
-
100
I
•
600
500
400
II
-
~
300
300
250
2
200
2 00
,
I
-- ~~
tqJ~. .-
~
IIIII11
t-..l
,.
100
-
IMmiI 800
900
'.
I
ahr,f-h----
-- Ffttl
--
1000
73
, I00
TRUE CRITICAL PRESSURE OF HYDROCARBON MIXTURES
-
..
1.18
!l.0
1.20
-
-
4.0
1.2.4
1.2.2.
-
1.2.6
M
4.0
-
3.0
2.5
2.5 DETERMINED BY MULTIPLYING ITS PSEUDO-CRITICAL
2.0
PRESSURE BY THE RATIO OF TRUE TO PSEUDO, CRITICAL PRESSURE, PTC/Ppc. THIS RATIO IS GIVEH BY THE CURVE AS A FUNCTIOH OF THE RATIO OF TRUE TO PSEUDO CRITICAL TEMPERATURE. TTCl1l'l;.l
.
""1l1i"
t:J+l '.5 H-t-t+t-H-t++t-H-t+'
J
mI.5
~ REFERENCE:
1.02
2.0
SMITH AND WATSON.
1.08
1.04
74
IND.
ENG.
CHEM,
29.
140B (1937)
Section 7
THERMAL PROPERTIES Specific Heat Since hydrocarbon vapors deviate considerably from a perfect gas, except at low pressures, their specific heats arc a function of pressure as well as tempera-
ture. However, vapor specific heats at higher pressures have limited application as enthalpy correlations may be more readily used for thermal calculations. For this reason, the specific heat chart~ for gases and vapors (pages 88 to 91) arc given only for low pressuros (0-1 atm) where deviations from a perfect gas are so small that specific heat may be considered to be a function of temperature alone. The specific heat of a mixture of two or more gases at low pressures may be
calculated from either their weight fraclions multiplied by their specific their mole fractions by their molal heat capacities (MC.).
heat~
or
Two charts are given for the specific heat of. petroleum vapors, one on page 90 for crude fractions and another on page 91 of more general application to both pure hydrocarbons and petroleum fractions.' The chart for crude fractions is a modification of the Bahlke and Kay eorrelation 2 and the other the same type as a chart developed by Fallon and Watson. 3 Both ehart~ are believed to be somewhat morc accurate than the previous correlations and arc also representative
of additional data. The change in enUlalpy of hydrocarbon vapors with pressure at constant temperature may be calculated from the chart on page 92. While the ordinate refers to the difference in enthalpy from the vapor at infinite dilution, this may be construed as any low pressure (0-1 atm). This chart was used to compute the enthalpy of hydroearbon 4 and petroleum vapors at elevated pressures in the development of the enthalpy charts. Since the change in enthalpy at constant 1 These correlations {or petroleum fmctions nrc not quite consistent with the additive rule for mixtures. Since these curves apply directly to mixtures, the additive rule would hold only if the specific heals eit,1).cr were independent. of the liquid specific gravity or \'aricd lillcnrly with its reciprocal (directly with API). With neither of these conditions fulfilled, the petroleum vapor correlations have a fundamental inconsistency but the resulting errors are imperceptible as far as the data are concerned. 2 Bahlke and Kay, Ind. Eng. Chem. 21, 042 (1929). 3 Fallon and Watson, Nat. Petroleum News, 'l'echnical Section, R-372 (1944). 4. For the light hydrocarbons below hexane, there was a slight trend with molecular weight in the change of enthalpy with pressure at constant. temperature. This was taken into account by the use of other unpublished correlations by Gilliland (sce reference on the chart OD page 92) for these low-boiling hydrocarbons. 0
75
tottlGl"N"
A.L.)
CO?l A \1i'1B.·::1 \. .) PABLO M01.'"l'A
76
DATA BOOK ON HYDROCARBONS
temperature can be read directly from the latter charts, this generalized chart has little direct application but is included as one of the fundamental correlations. The chart for the specific heat of hyd!"Ocarbon liquids was developed" directly from liquid specific heat data on pure hydrocarbons and petroleum fractions. Since liquid specific heats were not used in the development of the enthalpy charts, this chart is independent of and not necessarily consistent with the latter correlations." For the sake of consistency, thc enthalpy charts usually will be used in preferencc to this spccific heat chart but, at the samc time, it is desirable to include an independent correlation of such a fundamental thermal property. Latent Heat of Vaporization The latent heat of vaporization of any compound is the din'erence in enthalpy between its saturated vapor and its saturated liquid at constant temperature and may be expressed either as a function of temperature or as a function of vapor pressure. The latent heats of low-boiling hydrocarbons and, also, higher-boiling normal paraffins of even boiling point arc plotted again t vapor pressure on pages 94 to 97. While the use of temperature instead of vapor pressure as the correlating variable would have advantages, it would also result in the curves crossing each other, thus making the plots difficult to reae!. The latent heat charts were derived by using a direct proportionality between the molal heats of vaporization of any two hydrocarbons at the same reduced pressures.' For the lower boiling hydrocarbons, the latent heat data were smoothed out and extrapolated by the use of a reference compound (ethane, butane, or hexane). Where no data were available, as in the case of a few of the light hydrocarbons and all of the higher-boiling normal paraffins, the latent heats were calculated directly from this reduced pressure relationship. The slope or proportionality constant was predicted from the normal boiling point of the hydrocarbon. The latent heat of vaporization of other hydrocarbons may be calculated from the normal paraffin cun'es by the usc of this same relation. That is, the unknown compound will have the same molal heat of vaporization as a paraffin of the same normal boiling point at the same reduced pressure. In the case of petroleulll fractions, the mean average boiling point is used for the normal boiling point and the reduced pres ure i computed f!"Om the pseudo-critical pressure of the mixture. The "vapor pressure" of thc fraction corresponds to that of a pure • A modificatIOn ot a correlatIOn oy Tne M. W. Kellogg Co., New York, N.Y. • The enthalpy eharls were derived from: (J) the vapor specific heaL eo",elnlions (0-1 aIm); (2) lhe generalized chart for change of enthalpy with pressure; and (3) the latent heat relations. Inasmuch as the inaccuracies of all three correlations accumulate in Lhe jjqu;:! enthalpics 01' specific heals, the agreement wit.h the liquid specific heat chart may be ~nnsidcrcd quite good as average deviations between the two are around ±3% wlt.h a maxImum of about 6%. 7 Maxwell, Ind. Eng. Chem. 24, 502 (1932).
THERMAL PROPERTIES
77
hydrocarbon of the same normal boiling point at tbe temperature of the fraction and ""ver relers to the bubble point, dew pofnt, or operating pressure 01 the system. Since the difference in enthalpy between the liquid and the saturated vapor of a pctroleum fraction always involves change of enthalpy of the vapor at constant temperature in addition to latent heat, except at low pressures, the enthalpy correlations are much more convenient to use than these individual therm,.l properties. Thc following examples illustrate the usc of the latent heat charts: Example 1. Compute the latent heat of benzene at 1 atm.
The boiling point of benzene is 176.2°F and its critical pressure is 47.9 atm. The molceular weight of a normal paraffin boiling at 176.2°F is 91.5 and its critical pressure 28.3 atm. The vapor pressure of the normal paraffin corresponding to a reduccd pressure of 1/47.9 ( - 0.0209) is 0.0209 X 28,3 - 0.59 atm. The molal heat of vaporization of the normal paraffin at 0.59 atm is 91." X (146 BTU/lb) - 13,360 BTU/mole. The latent heat of benzene at 1 atm is then equal to 13,360 BTU/malo or 171 BTU/lb. The Bureau 01 Standards Circular C461 gives 169.3 BTU/Ii> '.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:<0 X 460 39.7
~
-
691°R <> 231°F
For the mixture, the work of compression = 0.500 X 1354 = 1482 BTU/mole <> 8f! BTU/Ib
+ 0.500
X 1610
The final temperature of the mixture is assumed to be the temperature I, at
which 0.500[Ji Jre (t, 7.91 atm) - Jilfc(l1l°F, 7.91 atm)J ~
0.500[Ji Il (23I°F) - flll(t)] = 0.500 X 6.97(231 - t)
Since it is necessary to use enthalpy for evaluating UTi lie, t will be determined by trial and error. Assume t = 140°F. Interpolating between the charts on pages 99 and 100, 0.500 X 34.2[3·12 - 328J = 0.500 X 6.971231 - 140] 240 "" 317 Assume t = 148°F. 0.500 X 34.2[346 - 328J = 0.500 X 6.971231 - 148] :l08 "" 290 By interpolation, the final temperature is I/.7°F. While the foregoing procedure permits the ~follicr diagrams to be used for mixtures of hydrocarbons and inert gases, the method of combining the cnthalpics and the temperatures of the two portions of the mixture is theoretically incorrect. In this procedure it is assumed that if two gases, ha,-ing different thermal propcries, arc compressed individually from the same initial temperature and pressure
86
DATA BOOK ON HYDROCARBONS
to the same final pressure and then mixed, the resulting thermodynamic properties of the mixture will be the same as if the gases were mixed initially and then compressed. This assumption is not quite correct and will always lead to small positive errors in the work of compression and temperature rise. The errors usually will not exceed a couple of percent with a maximum of about 50/0 if the average molecular weight of thc hydrocarbons is not greater than 50 and the compression ratio is not greater than 5: 1. As an alternative to this method, the equations for an ideal gas may be applied to the entire mixture, provided the gas law correction facwr for the hydrocarbons, P-YC, is not less than 0.05. In arriving at an average molal specific heat at constant pressure for the mixture, the molal specific heats of the individual components at 0-1 atm should be used irrespective of the initial and final pressures of the compression. The following equations apply to this alternative method:
= (y HCIlIlC + Yalla + Ybllb + ... ) (MCP)m = YlIc(MC p )llc + y.(MCp). + Yb(MCph Il ..
K =
AU
Tz
where
Il
+ ...
(MC p )., (MCp) .. - 1.99
= K K-
1
1l.,R1"
[(7I'z)K;: 1 71'1
]
1
= (::) \-1 T 1
= correction factor for deviation from the ideal gas law at initial conditions
Y = mole fraction of any component MC p = molal specific heat at constant pressure (0-1 atm) and at the average temperature MC p K=MCv d' b t' compression = work of AH = change in enthalpy during an a la a IC compression . 0 7", Pz = initial and final temperatures m R 71' = initial and final pressures 71'1'UC = subscript referring to the ~otal hydrocarbons = subscripts referring to mdlvldual mert gases a, bJ et c. Example 7 will be recalculated by the alternative method:
For the hydrocarbon portion: = 1.91 = 0.041; 1l1IC 460 . T r = -575 = 0.80e,
= 0.966
THERMAL PROPERTIES
87
For the hydrogen: I' =
1.000;
MOp = 6.97
For the mixture: I'm
MOp K
= 0.500 X 0.966 = 0.500 X 1.000 = 0.983
+ 0.500 X 6.97 =
=
0.500 X 14.0
=
10.5 _ • 10.5 - 1.99 - 1.230
10.5
1 235 [( 6 ) •. 235 6H = 9:235 X 0.983 X 1.99 X 460 1 \7 1.235 39
-
1]
~
4740[1.311 - 1] = 1470 BTU/mole as compared with 1482 BTU/mole previously calculated. T, = 1.311 X 460 = 603°R "" 143°F as compared with 147°F by the first method. IC desired, this alternative method may also be applied to hydrocarbon mixtures if I' at tbe initial conditions is ?!at less than 0.95. GENERAL REFERENCES Communication from The M. 'V. }{ellogg CO' l New York, N.Y.
Gary, Rubin and Ward, 11ld. E1lg. Chem. 25, 178 (1933). Gilliland l Unpublished data, Mass. Inst. Tech.
Keenan and Keys. "Thermodynamic Properties of Steam1 " John \Viley &; SODS (1936). Misc. Publicatio1l of Bur. Standards, No. 97 (1929). Nat. Bm. Statuiards Circular C461 (1947). Sage, Webster a-od Lacey, Ind. E1lg. ahem. 29, 1309 (1937). Weir and Eaton, Ind. E1lg. Chem. 24, 211 (1932).
'i
400 _, -~'l ~-;r,TIE' f fiffi'·[ili' .:;, .. .
20 ~-
~.
-
., .<+;;
600
800
1000
1200
-
.~
SPECIFIC HEAT OF MISCELLANEOUS GASES
" -,j~,tI1~ti
0-1 ATMOSPHERES
'1-···.. n.
0.26
..
·"l.IP
0.26
'
·f,.I~:;t+
.,.,
I£l~il:!~ill :
.1 .... •
gJ
_:.li~;:tI~JtIni:t _ti •• lf:t;.t±t:-:-~
f
l~n'd
0.60 c",
Ii
;", "·,'tT'
",,~-. ,..... t· :..:' ,'..
0.24
~TI!H~!"_,
I"
·1' "' ;:t'
eI
I
l!fi !;rei iJ:j±t'.:±i'lt!#t~!· 'l:~ ;-t:rTti-r"; H~l~
!i
~I":t
•
.. ..... ~,
;
0.22
0,20IlU'ill"~'
In":'Tl7'i""
,'", 11,,11
tU" 'I 200
",
--'
400
NAT. BUR, STDS, CIRCULAR C461 (1947) KEENAN AND KEYES. 'THEfNODYNAMIC PROPERTIES OF STEAM' JOHN WILEY AND SONS ~O.46 ( 1936)
800
1000
1200
1.2
1,0
. . . . .9
.8
..6
.4
300
"200
89
o
+f~-l.J..Ht+Ji
no IJij
'-.r;
SPECIFIC HEAT OF PETROLEUM VAPORS
CRUDE FRACTIONS AND FOR OTHER PETROLEUM FRACTIONS THE FOLLOWING MULTIPLYING FACTORS SHOULD BE USED:
.8.
• I •
CRUDE FRACTIONS. 0-1 ATMOSPHERES
J
.9B .96
i
I,
I
I
.7
r
lit
'I
ffiiI
rn1
THIS CHART APPLIES TO PARAFFIN BASE
MIXED BASE CRUDE FRACTIONS ASPHAlT" to ..
[1fI" ! j
'UlIJ
HiIrm; IiiI~I r; 1lIlftfi nm9
~
1tI111 Ufl IIIIHII11H4II$! "T
-, I ,ql
I
I
1
·,11111
~ , "
.7
tm!
"
8
II
TIfflI
OOB
.6
.fJi.llum·H1 HillllID~B;'lit
I"
n" I
iElHI
r,
.5
"
i
~'
; l.1oof1
~likuiJ ~""
IlIf
II
1m
_
r: ,.... t· ~ ,_._._.t. ;.. II.t +
I - _
•
"·hi ,~.
~
100
.; I! f'l.l·
iTt
.,·1
Hl:IlW @h!±Hil rl :4
tr~ii:b
:W.l.I±! .'
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700
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MOLLIER DIAGRAM FOR ETHYLENE
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440
420
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t-; -
360
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130
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540
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440
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360
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340
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310
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.90
Section 8
DENSITY Low-Boiling Hydrocarbons The specific gravity of the saturated liquid, from low temperatures to the critical point, is given on pages 140 to 142 for a number of low-boiling hydrocarbons. A hydrocarbon mixture is assumed to be an ideal solution, and its specific gravity can be calculated by adding the products of the specific gravities of individual components times their volume fractions. This assumption is essentially true for members of a homologous series and is a good approximation for mixtures composed of hydrocarbons £l'om different series as long as no component is in the region of it-s critical temperature. Thermal Expansion of Liquid Petroleum Fractions The thermal expansions of liquid petroleum fractions at pressures up to 1500 psig were derived from the thermal expansion and compressibility corrclations of Watson, Nelson and Murphy.! As in the case of many physical propcrties of petroleum fractions, thermal expansion is more sensitive to averagc boiling point than it is to gravity, although both independent variables arc necessary to correlate the data properly. Up to 1.25 multiples of the volume at 60°F and 1 atm, it was found that gravity could be neglected and that the thenpal cxpansion could be represented by the molal average boiling point alone. Above this expansion of 1.25 volumes, gravity is introduced into the correlation in the form of characteri7,ation factor. For each average boiling point two lines are shown, one corresponding to a characterization factor of 12.0 and the other to 11.0. Interpolation and extrapolation may be made on the basis of characterization factor or, if preferred, gravity, which is also given for each curve. P-V-T Relations of Hydrocarbon Vapors A series of charts on pages 148 to 153 give I' ~ PV IRT, the correction factor to be applied to the ideal gas law for hydrocarbon and petroleum vapors. The correction factor is plotted as a function of reduced temperature, TIT" and reduced pressure, PIP" where 7' and P arc the temperature and pressure of the vapor and T c and Pro its critical temperature and pressure. A explained in the scction on Critical Properties, the pseudo-critical, not the true critical, tempcraturc and pressure should always be used for hydrocarbon mixtures. This method of using the pseudo-critical properties of the entire hydrocarbon mixture is not only more accurate but more readily used than the application of either Amagat's Law or Dalton' Law to the individual components. 10;/ and Cas Jonma/35, 85 (1936).
136
•
DE"SITY
137
Since there is c\'idcncc of ::omc lr('nel in p- with incr('a~c in molecular \\"eight for T,. ;> J .00, there are thrC'c Ect~ of charts for the rq:rinn where T,. i~ {.!;I'cntcr than 1.0, c-o\"('rinf.!: din'erent r:ln~c~ of ll1ol<.'('u]:lr ""('i,,rla. Bt'!o\\" 1',- = 1.0, Ihe data arc jn~llmcjcnllo take into account n similar trend, so a ~illglc chart 2 covers the entire molecular \\'eight range. if olher gases (1-1,,0,,1-1,0, clc.) arc present in a mixlUl'e 01 hydrocarbon
Yapors, an cfTccti\'c Jlrc~surc equal tn 7rV?!IIC E"hould be used to obtain the rceluted pressure of the h~'drocarbon portion. Likcwi:::e, if it. ncc(':;:sary to lake int.o account ~as Jaw dc\'iations for any of the oth,,1' gn::=c<:, p. should be determined for carh of thc[o;(' g-ascs at an cO'celi\"(, Jln'.~:'llrc C''ll,lal to the total pressure multiplied by the square root of its muIr frnction. The molal \'oIUine is then calculated by Amilgnr~ Law, 111' V'" = - (y + y,,,, + y,,,, + ... )
is
11'
"C""C
\\-here I'm is the molal \'0Iu1l1e 01 the mixture, lhe subscript He relers lo lhe lolal hydrocarbon fraction, and the subscripts 0, b, et.c., rcfer to other gases. l.:'sua.Jly P-a/ jJ.ln etc., mar be taken as l.OO \\;ith ycry little error, sincc most of these gases approximate a perfect ~as at thc efl"ecti\'c pressures encountered. In the absence of other data, the hydrocarbon charls may be u,ed lor these gases.
VAl..UES
OF
GAS
Pressure
Lhlsq in. abs Lb/sq It abs ALm
Alm Atm ;\1 m or Ilg Lb sq rl abs
Cli It Ih-mole Cu It/I h-mole Cu It/lh-mole Liters, g-Illolc Cu III I h-mole Liters ~-molc Cu It I"-mole
CONST..... ~T--R
Temp.
n
'n
10.73 1545 0.7302 0.08205
OR
'H
OK
0J{
OK OK
GENERAL REFERENCES l1e:l.ttie, IIar/lork :lIlrJ Poffenberger, J. Chem. Phys. 3, 93 (1935). Beattie, Ka.\' and I\:aminsky, J. Am. Chem. Soc. 59, lii~9 (I93i). l1entlie, Rim:lnl and ~u, J. Am. Chem. SQr. 61, 2li (103D). Tnlcl'n:llional Criticnl Tnhl('f:, Vol. lIT. Kay, /11'/. "ng. Chem. 28, IOI~ (1030); 30, ·1,,0 (J03S); 32, 3,;8 (l9~O). J(cl~o:lnd FC'I!olinl!;. J. Am. Chem. Sn 62/ 3132 (19-10). L(>\\'i~, Inl/. Htl(l. ('hem. 28, 2.37 (I!l36). S.a~c :md L'l(,cy, Ind. Hllg. Choll. 30, (l7:~ (ID3R). Sal-!C'. Srh:t;lr"ma :,"d Lacey. hid. En9. ('Iu'm. 26. 121S (1!l~-I). S:l~C, \\'C"h<.;tC'1' :inti Larcy. I nIl. [.,'lIq. r1l,.,.". 29, G;j.". II~S (193i). Smith, I3c:\ttic~nd K3~·. J. Am. Chl"-'. Soc. 59,1:')1;)7 (lfl:li). c
,.
;!
COpC'. Lewis and Wcber, Ind. P;,IO. Chem. 23, 88i (931).
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Seclion 9 VISCOSITY Absolute Viscosity ]n the metric system the unit of viscosity is the poise which is equivalent to
B
force of I dyne pcr sq cm shcaring a liquid at the ratc of I cm pcr sec per cm. By reduclion to minimum dimensionality, the poise becomes I g/(cm) (sec). The corrcsponding English unit is I Ibl (ft) (sec), or (pouodal) (sec) I (ft) 2, which is equal t.o 14.88 poises. Howcvcr, the unit of viscosity most commonly used is the centipoisc (0.01 poise), which happens t.o be the viscosity of water at almost cxactly 68°F. Thereforc, the absolute viscosity of any fluid in centipoises may bc considered to be oumerically equal to its viscosity relativc to water at 68°F. Kinematic Viscosity
Since thc density of thc liquid involved in the measurement of viscosity by the standard industrial viscometers, it is necessary to introduce kinematic viscosity,
which is the absolutc viscosity of a fluid divided by its density at the temperature under consideration. The metric units of kinematic viscosity corresponding to
poises and centipoises arc stokes and ccntistokes, of which the latter is more commonly used. The kinematic viscosity of water is I centistoke at just about 68°F. Industrial Viscometers
The industrial viscometers which are widely used throughout the petroleum industry in this country are the Saybolt Thermo for rcfined oils, the Saybolt Universal fur lubricating and gas oils, and the Saybolt Furol for crude residua and heavy fucl oils. The Redwood Standard and Engler viscomctcrs arc used mostly abroad. Curves for the conversion of these standard viscometer measurements to kinematic viscosity arc gi\'cn on pages 158 to 160.
Except for the EnglCl' instrument, these conversions arc slightly affected by the tcmperature at which the viscosity is mcasured, but this effect has been neglcctcd in thc prescnt convcrsion charts, While Saybolt Universal viscosity may '>e measured at anyone of several temperatures, lOO°F, 130°F, or 2lOoF, the maximum variation between the temperature extremes in the conversion to kinematic "is('o~ily is only 3% and, above kinematic viscosities of 5 crntislokes,
it is Icss than 270. The va,'iation between the extremes of the Redwood Standard instrument (70°F to 200°F) is appreciable at low viscosities but does not exceed 3% above 10 centistokes. Saybo!t Thermo viscosity is normally measured at room 155
156
DATA BOOK ON HYDROCARBONS
temperature and Say bolt Furol at 122°F so that it is usually unnecessary to consider conversions at any other temperatures for thcsc instruments. Change of Viscosity with Temperature Viscosity-temperature curvcs arc given for pure hydrocllrbons and crude fractions on pages 161 to 165. In the absence of other data, these curves may be used to approximate viscosity-temperature relations for othcr hydrocarbons and petroleum fractions if the viscosity is known at only one temperature. However, if the viscosity is known at two or more temperatures, the charts on pages 166 and 167 should be used for linear interpolation and extrapolation. Viscosity Index Viscosity index is a generally accepted criterion for evaluating lubricating oils with respect to change of viscosity with temperature. The vi cosity index of any oil may be read directly from the charts on pages 168 to 172 if its viscositics at 100°F and 210°F arc known. If these particular viscosities are not available, but viscosities are known for two other temperatures, the viscosity-temperature charts on pages 166 and 167 may be used to find the values at 100°F and 210°F. Viscosity Blending To predict the viscosity of a blend of two or more fractions at any given temperature, the blending index for each fraction is determincd from its viscosity at this temperature, using the chart on page 173. The blending indexes of the individual fractions are additive by volume fraction and the resulting sum may be converted to the viscosity of the mixture by referring to the blending chart again. If the viscosity of one or more of the components is not available at the desired temperature, it must be converted.to this temperaturc, since blending indexes arc additive only at constant temperature. The viscosity of a blend of two stocks may also be obtained graphically by using the viscosity-temperature charts. A straight line eonnccting the viscosity of the lcss viscous stock on the 0°1" abscissa and the more viscous stock on the 100°F abscissa reprcsents the locus of the viscosity of all blends of these stocks. The ordinate corresponding to the percentage of the more viscous stock-whcreby the temperaturcs between 0°1" and 100°F are considcred per~entages-represents the viscosity of the blend. While thc blending indcx chart was derived from the ordinate scales of the viscosity-temperature charts, the two methods will differ slightly since thc tcmperature divisions vary between 0°1" and 100°F. Viscosity of Gases While pressure has very little effect on the viscosity of liquids except near the critical temperature, its effect on gases may be considerable, especially above the critical pressure. The change in viscosity of a gas or vapor with p1'essurc
,
VISCOSITY
157
may be predicted from the chart on page 177. By the use of reduced temperature and pressure, this chart provides a generalized correlation of the ratio of \'iscosity at any temperature and pressure to the viscosity at the same temperature and
atmospheric pressure. The viscosity of a mixture of two or more gases at atmospheric pressure may be computed by the following formula:
Z
m -
N,Z,v'M; + N 2 Z2VM; + ... + NnZ.VM: -'---'------==----=-'-="'--''------'-----~==_----'' N,VM; + N 2VM; + ... + N.-vM n
where Z... = the viscosity of the mixture N" N 2 , etc. = the mole fractions or moles of individual components Z" Z2, etc. = the viscosities of the individual components M M 2 , etc. = the molecular weights of the iudividual components
"
The chart for change in viscosity with pressure may be applied to mixtures by using the pseudo-critical properties of the mixture to determine reduccd tempera-
ture and pressure. GENERAL REFERENCES
ASTM Standard Viscosity-Temperature Chart. for Liquid Petroleum Products (D341-39), Charts C and D. Deale, "The Science of Petroleum," Vol. II, 1080, Oxrord University Press, New York,
N.Y. (1938). Comings and Egly, h,d. Eng. Che11l. 32, 714 (1940). Davis, Lapeyrouse and Dean, Oil Gas J. 30, No. 46, 92 (1932). Dean and Davis, Chem & Met. Eng. 36, 618 (1929). Edwards and Bonilla, !rId. Eng. Chem. 36, 1038 (1944). Etherington, Sc. D. Thesis, Mass. Inst. Tech. (1948). Evans, J. Insl. Pelroleum Tech. 24, 321 (1938). Forlch and Wilson, Ind. Eng. Chen!. 16, 789 (1924). I••ne nnd Dean, Ind. Eng. Chon. 16, 905 (1924). Lipkin, Duvi!'on and }(urtz, Ind. Eng. Chern. 34. 976 (19·12). Nat. Bur. Standards Circular CI,61 (1947). llnd Lacey, Ind. Bng. Chtm. 30, 829 (193 ). ~llge, Yale and I.,cey, l1ul. Bng. Chem. 31, 223 (1939). ~age
Watson, Wiell and i\furphy, Ind. EnO. Chem. 28, 605 (1936).
1000
10
20
60
80 100
300 400
600 800 1000
800
800
CONVERSION
600
500
TO
600 500
KINEMATIC VISCOSITY
400
400
300
200
100
60 -
50 40
-F_'':':
20
2000
10 _
1000
8
~
800
6
"
600 500
5
-
4
.. 400
k'.,:: ~#l ::: .... I::: :~. i
l~! :il~· J.".i , •.•
,h ' i1
~;
1
!lIT" I
400
!Hl·~
I
600 800 1000
15!3
2000
4000
I·'
300
100 10000
f,i
8000
200 -i.Jk.ti
600 800 1000
1000
.q
-=tot
800 600
6000
=11111
500 400 300
200
1000 -
100
::: cf
80 60 50
40 30 200
20
10
10 8
8.
6
6 :r;
~
5
4
4
3
3 >
t!
"
2
2
2
3
4
G
8
10
159
20
30
40
60
80 100
FROM SAYBOLT THERMO VISCOSITY
,.
4 SAYBOLT THERMO VISCOSITY -lOX (SAYBOLT THERMO TIME)
2
o
100
200
600
300
160
700
VISCOSITY OF NORMAL PARAFFINS
.1
-100
o
.I
100
200
161
300
400
1 1.8
...,.. _. -_.. ,-_r:r::
1.6 I
-:
.
+-
~±
VISCOSITY OF AROMATICS
:=T._.
1.4
-=1= .,
H· r -i' -r+--r-'. .-;...... -;- -
++-
-
r
I
,
..
± .
-
0.2
-
I
I I
0
50
..,...
~
.-i---
::!..
-
-
- I'-L
-
I-f
12
-
•
-- .
150
162
,.
-
- lJiIllI 100
-
.
-j I
,
L..j _
.
200
250
300
- -ri
40
KEY
,5:
10
-i:!...:?!=
r.l-
'i_: :;.'-=:'-~:. . ;
~~ ~ti ~~~£ ~~ .~ .~ 4' ; :1.si?=a" 6.0 ~i=LGl,{: -.M cFI..i.~:',r ;L,- l'!'=!o\2, ::EJ=L-" I-W::-;:;\ :jeil- _til 1-. \1 .,='1: -~l-+" = T:~ ,---+-+ - .-~-t-t \.tr--l=
80_. t 7.0 _0 __ e\'=~ •;~:il'ii ft
lc
1
I 2 3 4 5 6 7 8 9 10
:
26,8' API 33.2' API 35.0··API 36.~·API
38,2'API 39.2·API 43.~· API 46.0' API 48,7'API 51.soAPI
~ ~:..~:::~ ~~~=;~;:~
13
62.9·API
I
rt:'
"-+i-L
-i-tREFERENCE:
0.1
212 - 257°F.
NOTE- 801LING POINTS OBTAINED IN A HEMPEL COLUMN.
-
_'
RESIDUUM 527-572·F. AT 40"... 482-527·F. AT 40M", 437"462" AT 40"". 392-437'F AT 40"M. UP TO 39rF.AT 40""482 - 527·F. 437 -482·F. 392 -437·F. 347-392°F.
LANE AND DEAN. IND. ENG. CHEM. 16. 905 (1924)
1+
'i', . 100
200
!r~r~I' 300
163
400
500
600
700
-;~
80 70 60
. •
...
=.=t=<="
VISCOSITY OF MID CONTINENT OILS
•
50
40
30.~
KEY I 2 3 4 5 6
20
7 8 9 10 II 12 13
~ .. .. =~---
l4~ 0 API ~3. °API o
RESIDUUM CYLINDER STOCK 23.l API HEAVY MOTOR OIL (lO5 SPGR) ROAO orL 24.2°API RED OIL o 26.l API LIGHT MOTOR O~. 27.1 °API LIGHT PARAFFIN OIL 32.8~PI
WHITE OIL
28.8°API 3D.OoAPI 35.2°API 35.6°API 40.4°API
I~
LIGHT PARAFFIN OIL PRESSED DiSTILLATE CRUDE otL MINERAL SEAL KEROSENE GASOLINE
15
GASOLINE
•
do' L
4
• -,.,
3
2
09,_I,~
...
_~c_o _ .• 5§". . . =
.=
0.4
:
..
0.3
_-
0-2
REFERENCE,
FORTCH AND WILSON. I
100
200
400
300
164
IND .. ENG. OlEM. II
II
II
500
I
II
16. 789 (1924). I
I
600
II
II
7'00
wo
~
90 80 70 60
c
~~...:-
-
_ _T";.
- VISCOSITY OF CALIFORNIA CRUDE FRACTIONS
50 40
~
30
20
,
KEY I 2 3 4 5
6
10
7 8
8 7
9 10 II 12
6 5 4
18.7 'API 20.3 'API 22.8 'API 25.0 'API 27.3 -API 30.4 'API 33.0 'API 38.2 'API 41.5 -API 44.3 'API 49.5 'API 54.5 'API
527-572'F. AT 482 - 527'F. AT 437- 482'F. AT 392-437'F. AT UP TO 392'F. AT 482-527'F. 437-482'F. 392 -437'F. 347- 392'F. 302- 347'F. 257-302'F. 212 -257'F.
40 MM. 40 MM. 40 MM. 40 MM. 40 MM.
NOTE- BOILING POINTS OBTAINED IN A HEMPEL COLUMN
3
2
REFERENCE:
0.1
100
LANE AND DEAN. IND. ENG, CHEM. 16. 905 (1924)
200
300
400
165
500
600
700
500
600
700
00
-:fl 10eo
VISCOSITY - TEMPERATURE CHART HYDROCARBONS AND PETROLEUM FRACTIONS KINEMATIC VISCOSITY. LOW RANGE
6.0 5.0 4.0
30 I I I I I I I I I I I I I I I II I I I II III II1II II i III1II1 \1IIII11 i 111I1 11111 11I11 11111 11111 II!! 111111111111111111111111111 II I 11 III ITTTRRTlI I I I I I I I I i I m 3.0 20 I I I ! I I I I I I ! I ! ! I I I I I I II II ! ! I ! ! I ! I I I ! I I " i I II ! III " III !!III 1!1111111 ! I I !I !!! !111111 111 !111 1!III!! I! !I!III !II I I I I I I ! ! I I I I ! ! I !
!!
! I I ! I I ! Ii! ! ! 12.0
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1.00
0.90
~
0.80
0.70~~
0.70
0.60~·
'060 I .
():
j:;:
o.sOt~
J"
!!
I
.0.50
Zi
s;2i
0.40
!
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0.40
0.30 I I I I I I I I I I I I i I I I I I I I II I I I II I I I i I : 1111 i II ; 1111 !l111 i 1111 ;11111111111111111111111111111111111111111 i 111111 i II I I I i I I I I 1I I i I I I I I i I I I I I I I I I I I 10.30 R~FERENCES:
A.S. tM. STANDARD VISCOSITY - TEMPERATURE CHARTS 0341-39 NAT. BUR. STD5. CIRCULAR C 461 (1947) WATSON, WIEN AND MURPHY, IND. ENG. tHEM. 28,605(1936)
0.20
o
o
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o
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!
,
!
,
!
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J
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......... 199 10,000,000
8
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0
200N I,.
1~9.
, 2~q"" ~
~
400
450
..
500
600
700
800
~
VISCOSITY - TEMPERATURE CHART
1,000,000 500,000 200.000 100,000 50,000 20,000 10,000 5,000
HYDROCARBONS AND PETROLEUM FRACTIONS KINEMATIC VISCOSITY, HIGH RANGE
,
10C\OOO 50,000 20.000 10,000 5,000
2,000
-
2,000
1,000 (/)
1,000
500 ~
~
~4
z
-... 0>
500
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!
,
kb!
, W'2ffi E' bt!-bikffi.::tt!"ttJ:tl'!ifl
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I
!
I
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w <.) I
>-
50~§
(/)
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IJIIEJ 20
z
i:
10
10 8.0
8.0 60 5.0
6.0 5.0
,
, ,
4.0 •
4.0 ,
,
I
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, ,
,
'
,
REFERENCES:
,
WATSON. WIEN AND MURPHY, IND. ENG. CHEM. 28,605 (193G)
, TEMPERATUflE.~ .·.F. ,
2.0
A.S. TM. STANDARD VI$CO$ITY- TEMPERATURE CHARTS 0341·39 NAT. BUR, STOS. CIRCULAR C 461 f'947)
3.0
, 34 33
VISCOSITY INDEX ALIGNMENT CHART 2.0 TO 50 CENTISTOKES AT 210°F.
32
140
31
130
30 29
120
28
I"~
27 26 25
110
100
-"S
24 23 22 21 20 19 18
90 u:
)(
Q
0
0
...
~
>-
0
w
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...
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.-
16 15
<:>
<:>
l/l
q
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17
t:
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60
;g $ 0
50
14 13
40
12 II
30
10 9
20
B
7
10
6 5
0
168
~50
440
140
VISCOSITY INDEX ALIGNMENT CHART
430 -
40 TO 60 SAY80LT SECONDS AT 210"F.
420
130
410
400 120
390
380
60
370
58
360
5;>
350
56
340
s.,
330
300
5,3
"'
~
290
~
280
'"0z
270 260 250
100
55
320 310
110
59
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230
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240
90
60
220 -
50
210 200 40
190 180 170
30
160 150 20
140 130 120
10
110 100 0
90
1G9
1650 1600
VISCOSITY INDEX
1550
140
ALIGNMENT CHART
50 TO 120 SAY80LT SECONDS AT 210°F. 130
1500 1450
120
1400 1350
110
1300 1250
100
1200 1150 1100 1050
90 u:
!<.' ° ~ C\i
°0 0
1000
Vl 0
900
<>
,,'"'"
0
800 750
Vl
~
0
m >
Vl
>-
in 0 <>
tJ:;
UI
....
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80
f--
t
z
850
~
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f-
950 -
x
w
0
70
'"
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60
~
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50
700 650
40
600 550
30
500 450
20
400 350
10
300 250
o
200 150 1.
170
3800
VISCOSITY
3700
INDEX ALIGNMENT CHART
80 TO 200 SAY80LT SECONDS AT 210"F.
3600
130
3500 3400t
120
::1
110
3000+t
100
2800:1-
90
3100 -
1900+
t
2700t 2600 i
u:
"0
x
Q
~
0
24COt ,
'"z 0 """
>!:::
T
2300+ 2200t 2100
f T
2000
0
0'" ":;'"
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70
60
0
al
>
50
'"
1900
80
"" :;;
2500!
40
1800 1700 1600 I T 1500+ ,
30
1400+
20
+
1300+
t +
10
1000:1-
0
1200+ 1100+
900t 800 1
171
I
7800 T 7600
VISCOSITY INDEX ALIGNMENT CHART
1400
100 TO 350 SAYBOLT SECONDS AT 210°F.
130
7200 120
1000 6800 6600 6400 6200 6000 5800 5600 5400 5200
u;
°
~
~
5000
0
z
0
<.)
4800 4600 4400
w
~
16>-
..
4200 4000 3800 3600 3400 3200
20
3000 2800
10
2600 2400
o
2200 2000 1800
172
0.2
0.3
0,4 0.5
2.0
35
30
80
25
60
20
40
~
15
10
10
20
200
300 400 500
2000
5000
10000
.030
.030
.028
.028
.026
.026
.024
.024
.022
.022
.020
.020
.018
.018
.016
.016
.014
.014 ;,
.012
;+
.012
~
.010
.010
.008
.008
.006
.006
.004
.004
.002
.002
o
200
400
600 174
100
80
~IIII-!!IIII 20
10
.BEALE.
THE SCIENCE OF PETROLEUM. VOL. 11. P.
. OXFORD UNIVERSITY
PRESS (1938)
0.1
I
I
o
100
200
300
175
400
500
o
200
mtmm
400
600
800
1200
1000
.040
.040
ABSOLUTE VISCOSITY OF .038
..038 .0.36 .034
.034. .032 .030
.032 .030
.•
.026 .026
.026
.024
024
.022 .020
.020
.
.018
,
"
•.018
.016
.014
.0/4
.012
012
, .010
.,
~
.008
.006
I
l' fr ~ f i r t•
o
200
600 TEMPERATURE -
176
800 0
F
iJ. .•
1000
.009
1+
;:j~ ~ 1200
.006
.3
.2
.4
.5
.6
.7
.8
.9
1.0
3.0
4.0 10
10 9.0 8.0
9.0
~~. ,~,E~E~,£,~~~~,~,~~,IN.~~.,.6:~"~, ,~~,;~~,~~~",r' " " "'~
8.0 7.0
7.0 6.0
6.0
I :: i
! I i I i i i I II II I VISCOSITY OF GASES AT HIGH PRESSURES
4.0
4.0
3.0
3.0
2.0
2.0
.., -.J
1.5
H;tt+tttf+
1.5
H , ,
"
11'i],11
1.0 T 11 .2
1.0
.3
.4
.5
.6
.7
.8
.9
1.0
1.5
2.0
3.0
4.0
5.0
6.0
7.0
80 9.0 10
Section 10
COMBUSTION Liquid Fuels The heats of combustion of fuel oils and petroleum fractions are expressed as a funcLion of gravity by the chart on page 180. Both the high and low heating values have becn correcLed for the average impurities other than water which are usually prcsent in oils of various gravities. These average impurities, tabulated on the chart, are fairly represcntative, although there may be appreciablc deviations for a given stock. In general, the heating valucs of average fuel oils arc within 1'i& of the curves. The heat available from the combustion at 60°F of liquid fuels is given on pagcs 186 to 188 for fuel oils of 5°, 10° and 15°API. Because of the small variation betwecn these charts, interpolation is unnecessary and the available heat at any tempcrature and percent excess air may be read from the chart which most nearly corresponds to the gravity of the fuel oil. If the impurities are known to be appreciably difTerent from the average values tabulated on page 180, the available heat may be corrected in direct proportion to the bydrocarbon portion of the fuel with sulfur considercd as inert material. Gaseous Fuels Heats of combustion of paraffin and olefin gases arc given as a function of molecular weight by the chart on page 181. The paraffin curves on this chart were used as a basis for deriving the charts on pages 184 and 185 for the heat available from the combustion at 60°F of dry refinery gases having high heating values of 1000 and 1600 BTU/S.C.F. Allo\\·ance was made for average impurities of 2.570 H 2 S and 2.5% inerts (equal parts CO 2 and air) by volume. As in the case of liquid fuels, the chart more nearly corresponding to the high heating value of the fuel gas may be used without interpolation with very little error. However, in correcting for variation in impurities, the available heat must be adj usted in proportion to the weight pm·cenlage of the hydrocarbon portion of the fuel gas. In making an adjustment for thc H 2 S content of thc gas, its volume percent may be distributed equally between the inerts and hydrocarbon portion as a good approximation. The following table gives relevant information for refinery fuel gases of 9.verage impurities: \78
COMI3 S
no"
Nominal HHV, BTU/S.C.F.-, .. _.•....
1000
1200
Wt. Percent Impurillc~ ... ~.
10.1
8.3
~p
0.60
0.73
16.5 1037
20.4 1248
G, of Fucl Gas ('-or - 1.0'-.
M.W. of H)'drocarbon Poruon. Actual HHV of He Purthm-U rU/ti.C.F. ' I
Calcubtcu by the perfcct gas hw at GOoF and
J
179 1400 7.1 0.86 24.3 1458
1600 6.2
a.utl 28.2
1609
1SOO
2000
5.4
4.00 1.2b
1.12 32.1 1879
36.1 2090
atm.
Properties of Flue Gas The CO 2 content of flue gas and the lI'eight ratio of flue gas to fuel are given both for liquid and gaseous fuels as a function of excess air on pages 189 <1nd 190. Since the effect of percent excess of air is almost imperceptible on the viscosity and thermal cunduetiyity uf flue gas, it has been ncglected entirely and each of these properties is expressed as a function of temperature alone.
;e . ,.it
~
tit~~'
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HEAT OF COMBUSTION OF FUEL OILS AND PETROLEUM FRACTIONS
,~
..
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_(,,!It. ::::
:;t;:t!1
".
-
IMPURITIES IN AVERAGE FUELS . " "A.P.I. I % S %INERTSI'M~Y,m'tV 20000
~
RESIDUAL FUEL OILS AND CRUDES o 2.95 1.15 4.10 5 2.35 100 3.35 10 I. 80 .95 2.75 15 I. 35 .85 2.20 20 1.00 .75 1.75 CRUDE OILS 25 .70 .70 1.40 30 .40 .65 1.10 35 .30 .60 .90
19500
.
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THESE VALUES REPRESENT AN AVERAGE OF CRACKED AND VIRGIN FUEL OIL DATA UP TO 20"API,AND THE CORRELATION ALLOWS FOR AVERAGE SULFUR AND INERTS (EXCLUDING WATER) FOUND IN AVERAGE FUELS. ABOVE 40"API THE CORRECTION FOR IMPURITIES IS NEGLIGIBLE AND THE CURVES REPRESENT PURE PETROLEUM LIQUIDS.
.
17000
o
10
20
30
180
40
50
60
19
4600
4409 4200 4000
001.
341
3200
2800 2600 2400 2200 2000
1800
1200 1000
10
181
900
ENTHALPY OF FLUE GAS COMPONENTS
, ~.
0-1 ATM.
B"· ~.
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;'If!- ~ ~ _:.:. -: t t
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1000
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200 . 1600
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2000
2200 183
2400
2600
2800
3000
HEAT AVAILABLE FROM THE COMBUSTION OF REFINERY GAS 1000 B.T.U.I CF (60· F)
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r. :.. j"
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2000 2200
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2000 2200 2400
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THE COMBUSTION OF
50 API FUEL OIL 15000
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188
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Section 11
FLOW OF FLUIDS Friction Factor
The friction factor for turbulent flow of all fluids (liquids and vapors) is expressed as a function of a modified Reynolds number (DUS/Z) by the chart on page 198 for both commercial pipes and smooth tubes. In the unstable flow region between values of DUS/Z of 0.135 and 0.390 (or approximately 1000 and 3000 in consistent units for DUp/p.) the turbulent flow curves have been extended to the stable streamline flow region. These extrapolated curves for turbulent flow give maximum values of the friction factor in the unstable region and are representative of the flow usually found in commercial pipes. For streamline flow the pressure drop may be computed directly from either of the formulas given on the chart, since the friction factor is incorporated in these formulas. Pressure Drop in Commercial Pipes To facilitate the determination of pressure drop for liquids in commercial pipe" the charts on pages 199 to 201 were derived from the friction factor curve and the formula for turbulent flow. The following example illustrates the application of these charts: Exampte I. Determine the pressure drop for 21,800 gnl/hr of gasoline f1owin~ through 800 ft of standard 6-in. pipe. The kinematic viscosity of the gasoline is 0.60 cs and its specific gravity is 0.750 at 100°F, which is the average temperature of the gasoline in the pipe.
21,800
.
Q/D = 6.065 = 3600 gal/hr/m. By following the dotted lines on the chart on page 200 as indicated for ~ 3600 gal/hr/in. to a kinematic viscosity of 0.60 cs, then over to the inside pipe diameter of 6.065 in., the value of AP/S ~ .38Ib/sq in. per 100 ft.
Q/D
The pressure drop for 800 ft of pipe will be:
t:.P
=
0.38 X 0.750 X
~:
=
2.3 Ib/sq in.
Equivalent Lengths of Fittings Data on the frictional resistance of fittings are usually correlated by the equation Ah ~ Ku 2 /2g,·where K is a constant for each type of fitting. However, 193
DATA BOOK ON HYDROCARBONS
194
in problems of fluid flow it is more conycnicnt to express thesc resistanccs as equivalent lengths of straight pipe for use in thc gcneral friction factor equation. Since the latter is a function of Reynolds numbcr "'hile J( i" an independcnt constant, it is nccessary to corrcct the equivalent lengths for variation in Reynolds number in inverse proportion to the friction factor. In the table on page 202 the equivalent lengths correspond to a Reynolds number of 10 and, for appreciably different values of the latter, should be multiplied by the correction factor on page 203. Example (Liquid Flow). Kerosene at lOO°F is being pumpcd at a ratc of 18,000 gal/hI' (gal/hr at 60°F) through 500 ft of standard steel 4 in. pipe in which there are eight standard elbows, one tee (side out) and two gate valves. Calculate the pressure drop through this line using the friction factor curve for the 1I0w through the pipe and the "K" factors for the fittings; check the result using the pressure drop charts and equivalent lengths for the fittings. The kerosene has an absolute viscosity of 1.5 cp at lOO°F, a specific gravity of 0.825 at 60°F, and a volumc cxpansion ratio of 1.025 at lOO°F relative to 60°F.
Q = 18,000 X 1.025 = 18,500 glll/hr at lOO°F U - 0.00680 X 18,500 - 78 f
-
(.1.026)2
-.
I
t sec
,.. 0.825 SpeCific GravIty at lOO°F = - - = 0.805 1.025 4.026 X 7.8 X 0.805 1.5
DUS ----z= ti
=
168;
f
=
0.0052
P (. ) _ 0.323 X 0.0052 X 0.805(7.8)2 X 500 _ 0 lbl . pipe 4.026 - 1.2 sq Ill. .
fi tiP (lttlllgS)
(8 X 0.45
=
+ 1 X 0.90 + 2 X 0.19)
(7.8)2 X 0.805 148.2
Ib/' 4.88 X (7.8)2 X 0.805 = 1.6 sq Ill. 148.2
= .
Total pressure drop = 10.2
+ 1.6
= 11.8 lblsq in.
Check
Uncorrected equiv. length of fittings = 8 X 6.6 Correction factor
(D~
S
=
16.8)
=
+ 13.2 + 2 X 2.8
1.1
Corrected equiv. lcngth of fittings = 1.1 X 71.6 = 79 ft
•
= 71.6 ft
FLOW OF FLUIDS Total equiv. length
=
500
Q
579 ft
18,500 .1.026 = 4600;
D=
tJ,:
+ 79 =
195
Z
S=
1.5
0.805 = 1.9
2.5 lb/tiq in./lOO ft
=
Total pressure drop = 2.5 X 0.805 X
~~~
=
11.7lb/sq in.
Example (Vapor Flow). Propane vapor at 90°F and an upstream pressure of 20 psig is flowing through 800 ft of 6 in. standard steel pipe at a rate of 25,000 Ib/hr. Determine the pressure drop through this line assuming the ideal gas law applies to propane under these conditions. At 90°F the viscosity of propane vapor is 0.0095 cpo Thc following cquation for isothermal flow of ideal gases and vapors can be derivcd by applying Bernoulli's theorcm to a diffcrential length of pipe and integrating thc rcsulting cquation between the limits, 0 and L: gI?T(P,' - 1'2 2)
1
U
=
J
2MP,2 [/L -
2m
I' J + In-' P 2
where U 1 = upstream velocity in ft/sec PI = upstream pre ure in )b/sq ft abs P 2 = downstream pressure in Ib/sq ft abs T = absolute temperatme-OF + 460 L = Icngth of pipe in ft m = hydraulic radius in ft = d/4 for pipes I = friction factor 9 = gravitatioual constant = 32.2 ft/sec/sec R = ideal gas law constant = 1545 M = molecular weight By substitution flnd rearrangement the abovc equation can be converted to a modified form of the equation for liquids, or
tJ,P = P - P = I
2
PI
+ In P24/P. + 1'2 [0.323 (If, D I
21'\
2
)
S U 1
I
'J
where PI, P 2 = upstream and downstream pressures in Ib/sq in. abs
S I = speci'fi c gravity . 0f l vapor re ' atlve D = pipe diameter in inches
I,0
M1' I water = 000 . 150 T
DATA BOOK ON HYDROCARBONS
196
Trial and error must be used in the solution of the above equation since P 2 is unknown. The friction factor, I, is independent of the variation of pressure since the mass velocity term, US, in the Reynolds number remains constant, U varying inversely and S directly with the pressure. D = 6.065 in.
S = 0.00150 X 44 X (20 90 + 460 Density
= 0.00416
X 62.4
+ 14.7)
= 0.00416
= 0.259 Ib/cu ft
U 25,000 X 144 X 4 _ 134 ft/sec , - 0.259 X 3600 X..- (6.065)2 DU,S,
Z
=
6.065 X 134 X 0.00416 0.0095 = 355; f = 0.0031
For the first trial assume P 2 Ii P =
= P,
0.323 X 0.0031 X 800 X 0.00416(134)2 6.065
=
9.
9 lb/
.
sq tn.
For the second trial assume P 2 = 24 Ib sq in. liP = 2 X 34.7 [0.323 X (0.0031 X 800 6.065 34.7 + 24 =
1.18[0.323 X-(0.409
+ 0.015)
+ 0.37) X 0.00416 X
(134)2]
24
X 74.5J
=
12.1 lb/sq in.
A third trial would give a liP of 11?,4 lb/sq in. In this example neither the initial velocity nor a contraction loss from a larger vcssel into the line was taken into account. If thc propane vapor wcrc flowing from a drum into the 6 in. line, it would be necessary to calculate an initial pressure drop as follows assuming isothermal flow: RT In P /P 01 M
= U,2
0.5U,2
2g+2g
The first term on the right-hand side is the velocity head, and the second term is the actual contraction loss due to friction. If the available head in the drum, Po, is 34.7 psia, PI is determined by trial and error and for the first trial U, is assumed to be 134 ft/sec. In Po/P,
=
1.5M 2 2RTg U,
=
-6
1.2 X 10
U,
In PO/PI = 1.2 X 10-6(134)" = 0.0216 PO/PI = 1.022 PI = 34.0 psia
•
2
FLOW OF FLUIDS
197
Since the differential is so small, 0.7 lb/sq in., a second trial is unnecessary. If this loss had been considered at the beginning of the example, the latter would then have been based on an upstream pressure of 34.0 instead of 34.7 psia.
GENERAL REFERENCES Beij, J. Research Nat. Bur Standards 21, 1 (1938). Chilton and Colburn, Ind. Eng. Chern. 26, 1183 (1934). Crane Company, "Flow 01 Fluids Through Valves, Fittings, and Pipe" (1942). Drew and Genereaux, Trans. Am. Inst. Chern. Eng. 32, 17 (1936). Foster, Trans. Am. Soc. Mech. Engrs. 42, 647 (1920). Gourley, Proc. Inst. Civil Eng., p. 297 (1910, Part 2). Karl' and Schultz, J. Am. Soc. Naval Engrs. 52, 239 (1940). Schader and Vanderlip, Cornell Univ. Eng. Exp. Sla. Bull. No. 130 (1935). Walker, Lewis, McAdams and Gilliland, "Principles 01 Chemical Engineering," pp. 71, 87-89, McGraw-Hill Book Co., New York, N.Y. (1937).
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.006
SDS
Z ~ ABSPLUTE VISCOSITY IN CENTIPQISES S ~ SPECIFIC GRAVITY RELATIVE TO WATER L ~ LENGTH OF PIPE IN FEET Q~ QUANTITY OF FLUID IN GAL.lHR. W' WEIGHT OF FLUID IN LBS.lHR. U ~. 006BO Q /0 2 ~.000816 W/SD 2
f!l-UI, I .,.,
ttT
, ; .;_i 1,,Ii ::1
.
05
WHERE Ll.P ~ PRESSURE DROP IN LBS.lSQ.IN. 0' PIPE DIAMETER IN INCHES
I
t eM
'1.495 X IO'5fSQ2L
D4
U ' LINEAR VELOCITY IN FT.lSEC.
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STREAMLINE FLOW
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60
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-C>P'PRESSURE DROP IN LBS./SQ.IN./IOOFT. D' ACTUAL INSIDE DIAMETER IN INCHES Q' QUANTITY OF FLUID IN GAL.lHR. Z' ABSOLUTE VISCOSITY IN CENTIPOISES S' SPECIFIC GRAVITY Z/& KINEMATiC VISCOSITY IN CENTISTOKES
li~rPRESSURE ,:n
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h-i-rll i dli II ill 11.11. i Ii ITTn1l11t11
r
DROP IN COMMERCIAL PIPES
rr- I .
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TURBULENT FLOW
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.
.:I.P= PRESSURE DROP IN LBS./SQ.IN./IOO FT. 0= ACTUAL INSIDE DIAMETER' IN INCHES Q: QUANTITY OF FLUID IN.GAL/HR. • Z: ABSOLUTE VISCOSITY IN CENTIPOISES S' SPECIFIC GRAVITY ZIS- KINEMATIC VISCOSITY IN CENTISTOKES
,
..
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.
PRESSURE DROP IN COMMERCIAL PIPES TURBULENT FLOW
STREAMLINE FLOW .:I.P'4.55 X 10-4 ZQ
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./lop, PRESSURE DROP IN LBS.lSQ.IN.lIODFT. 0' ACTUAL INSIDE DIAMETER IN INCHES Q'QUANTITY OF FLUID IN GAL.lHR. '. Z 'ABSOLUTE VISCOSITY IN CENTiPOISES S'SPECIFIC GRAVITY Z!S'KINEMATIC VISCOSITY IN CENTiSTOKES
.' +:- ......-r.
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PRESSURE DROP IN COMMERCIAL PIPES TURBULENT FLOW
~I'
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STREAMLINE FLOW AP'4.55X 10-4ZQ
04
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10000
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60 00 100
EQUIVALENT LENGTHS OF FITTINGS Pipe size -
Equivalent Lengths' - Feet
Incbes
Elbows
Vah'es
1.D.
Tees
Dends
Nominal
Inside Diam.
0.0.
Standard
Extra
Strong
Globe! Galc
-- -.WKt 10
Stanrlard·
Long 00' Sweep RID = 6
----
45°
= 1.5
Close Return
= 1
Side Out
End Out
Run of Standard
.25
.21
.75
.no
1.3
.30
0.4 07 o.n 1.1
03 0.6 0.7 1.0
0.3 0.5 0.6 0.8
1.0 1.7 2.3 3.9
1.2 2.0 2.7 3.4
1.7 3.0 3.9 5.0
0.4 0.7 0.9 1.1
2.6 3.4 5.0 6.6
1.8 2.3 3..4 4.4
1.5 1.9 2.8 3.7
1.2 1.6 2.4 3.1
4.4 5.7 8.4 11.1
5.3 6.8 10.1 13.2
7.7 9.8 14.6 19.1
1.8 2.3 3.4 4.4
29. 38. 48.
3
.45
.30
0.3 0.4 0.6 0.7
4.0 6.8 n.o II .5
0.6 1.0 1.3 1.7
17.6 23. 34. 44.
1
0.540 0.8·10 1.050 1.315
0.364 0.622 0.821 1.04n
0.302 0.516 0.742 0.n57
1)1 2 3 4
1 !l00 2 375 3.500 4.500
1.610 2.067 3.068 4.026
1.500 1. 93!) 2.!l00 3.826
59. 75. ll2. 147.
1.1 1.4 2.1 2.8
6 8 10 12
6 62.5 8 625 10.75 12.75
6.065 7.981 10.020 12.000
5.761 7.625 9.750 II. 750
220. 290. 360. 440.
4.2 5.5 7.0 8.3
66. S7. llO. 131.
10.0 13.1 16.5 19.7
6.6 8.7 11.0 13.1
5.5 7.3 9.1 10.9
4.7 6.1 7.7 9.2
16.6 22. 27. 33.
1!l.9 26. 33. 40.
14
14.00 16.00 18.00 20.00
13.25 15.25 17.18 19.18
-
480. 560. 630. 700.
9.2 10.6 11.9 13.3
145. 167. 188. 210.
22. 25. 28. 32.
14.5 16.7 18.S 21.
12.0 13.n 15.6 17.5
10.2 11.7 13.2 11.7
3(; . 42. 47. 53.
44. 50.
63.
56.
82. 91.
~
)1
%
""o ""
13.3 23. 30. 38.
Angle
16 lS
20
-
• The equivalent lengths tabulated correspond to a value of on the opposi te page. XC;, S KC;'S pP=-X- =-2q 2.31 148.2 : For swing che('k valve, usc ~~
or globe
D~S
valve equivalent lengths.
J)[''i
=
63.
57.
72.
6.6 8.7 11.0 13.1 14.5 16.7 18.8 21.
10. For other values of - Z:"" , apply correction factor from the chart
o
CD
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0
'"
FRICTION LOSS DUE TO SUDDEN CONTRACTION AND ENLARGEMENT TURBULENT FLOW IN PIPES
LOSS DUE TO CONTRACTION
LOSS DUE TO ENLARGEMENT
(U2)2 Fe' K64:3
F
E·
(U,-U2)2 64.3
A2' F • K • U, • Ue.
DOWNSTREAM AREA FRICTION LOSS, FT. OF LIOUID FACTOR FROM CHART UPSTREAM LINEAR VELOCITY - FT. / SEC. DOWNSTREAM LINEAR VELOCITY • FT. / SEC. lip· PRESSURE DROP DUE TO FRICTION LBS./SO.IN. LOSS S • SPECIFIC GRAVITY OF FLUID AT TEMPERATURE UNOER CONSIDERATION
.4
.3. .~
.2
.~-
i
.
:±±: :.T'/.
,.
•1
-'- .
.
. WALKER. LEWIS. MC ADAMS ~D GILLILA~D. 'PRINCIPLES or CHEMICAL ENGINEERING. PP. 87·89. MC GRAW·HILL BOOK CO. (1937) u. r. • . .•. •• l-!r.. 0
,
y..l~:'
:t '-I: .1
.2
.3
.4
.5
.6
.7
204
.8
.9
1.0
o
_7
2.
DISCHARGE CHARACTERISTICS OF RECTANGULAR AND CIRCULAR WEIRS 3400 6000 1
'
3200 "
3000
2400 2200 .2000 1800 1600 1400 1200 1000 800 600 400 200
7 0
0
205
.01 04
.02
.3
.45.6
.6
I
3
2
.03 .02
610..04 •
REFERENCE,
COLBURN.
IND.
ENG.
CHEM.
26.
.03
1183 (1934)
.01
01
PRESSURE DROP ACROSS TUBE BANKS
.006 .005
Ii .....L.,. N""('- ')2'(~)
. 004
liP' 3B90
S
GM
Os
•
.008
006 005
f'
.003 "")' FOR TUBE ANO SHELL HEAT MULTIPLY 4P BY A BUNDLE FOLLOWS; 0.'0 FOR SQUARE-TUBES 0.40 FOR SQUARE - TUBES
EXCHANGERS FACTOR AS
.002 IN LINES
AT 48·
GM IS EVALUATED AT CENTER ROW OF TUBES
.01
_
..001
.006 .006 .004 .003
__
"
.002
'
.001
.001 .0006 .0006 .0005 .0004 .0003 .0002
10
006 liP
:: PRESSURE
0006 0005
,DROP ~ LB./SO. IN.
N 'NUMBER OF ROWS OF TUBES
S GM
Do
GRAVITY OF FLUID RfLATIVE TO WATER t
0004
,MAXIMUM MASS VELOCITY' LB'/SEC./SQ. FT. (THROUGH MINIMUM FREE CROSS-SECT. AREAl,
.0003
:: SPECIFIC
1.
:: OUTSIDE TUBE DIAMETER· INCHES 1~ ::; MINIMUM CLEARANCE BETWEEN TUBES-INCHES;
Os f' 'FRiCTION FACTOR FUNCTION ZF -; FILM VIS'COSllY - CENTlP
20
30 40 5060 60 100
0002
200
300 400
206
600 6001000
Section 12
FLOW OF HEAT
18
6
5
4
3
2
Heat Transfer The film transfer cocfficient for liquids flowing inside tubes (page 211) is based on the Sieder and Tatc corrclation l which is generally acceptcd as the most rcliable for this type of hcat transfcr. Thc chart on pagc 212 for the outside film coefficien~ for flow across tube bundlcs was derived from a corrclation by Chilton and Colburn 2 with thc consistent units in the dimcnsionless terms replaccd by more common units. Comparison of limited data with this correlation has indicated that the film coefficient should be multiplied by the "bundle factors" given on the chart when GM is taken as the mass velocity at the center row of tubes. Kon-uniformity of flow and by-passing bctween the tube bundle and shell appear to bc the principal reasons for this diffcrence. Thermal Conductivity of Petroleum Liquids
)8
)6
)5 )4
)3 )2
Attempts to correlate thermal conductivity of petroleum liquids as a function of gravity in addition to temperaturc havc resulted in contradictol'y trends with °API gravity.3.4 In view of this inconsistency and since Smith 5 has shown that, at 86°F, a single value represents the reliable data belter than either trend with gravity, the relabon for the thermal conductivity of petroleum fractions on page 213 is shown as a function of temperature alone. This chart may also be used for pure hydrocarbons, although the data on low-boiling aromatics arc about 1070 higher than the curve. Thermal Conductivity of Hydrocarbon Gases As most data on the thermal eondueti"ity of hydrocarbon gases were obtained at room temperature, it is was necessary to find some means of extrapolation to higher temperatures. This was done by using two different methods: (1) assumption that the Prandtl number is a constant independent of temperature and (2) employment of Sutherland's equation. As the results of the two methods became more divergent with increasing temperature, it was a question of selecting either one or the other or using an average of the two. An average was chosen I Sieder and Tate, l1ul. Eng. Chem. 28, 1429 (1936). 2 Chiltonllnd Colburn, Ind. Eng. Chem. 26, 1183 (l934). 3 MUic. Publication 01 Bur. Sta>l(lards, No. 97, 24 (l929). 4 Kaye and Higgins, Proc. Royal Soc. 117, 459 (1928). 5 Smith, Trans. Am. Soc. Mech. Engrs. 68, 719 (l936).
207
DATA BOOK ON HYDROCARBONS
208
since, while it was felt that the Prandtl number was probably more relia~le, the Sutherland equation gave lower and consequently more conservative values. In view of the uncertainties of these extrapolations any refinement beyond the use of a straight line was unwarranted. Consequently, the chart on page 215 gives the thermal conductivity of hydrocarbon vapors as a linear function of temperature for various molecular weights. Logarithmic Mean Temperature Difference In the transfer of heat between two fluids, the log mean temperature difference applies to flow that is either entirely countercurrent or entirely concurrent. nder conditions where there is a combination of these two types of flow, such as a heat exchanger with more tube passes than shell passes, Nagle 6 has shown that a correction factor should be applied to the log mean temperature difference. This correction factor is given herein by either one of two types of charts, the first on page 218 and the second on pages 219 to 221. The chart on page 218 may be more convenient to use when the factors R and A do not approach unity. If these factors arc near to unity, it is necessary to use the other charts. The following example illustrates the application of these charts: Example 1. Determine the correct temperature difference and the number of shell passes required in the heat transfer between two fluids having the following inlet and outlet temperatures:
Shell side: T I (inlet) Tube side: t l (inlet)
R = TI
= = -
t2 -
m =
t2 TI
-
400°F; 275°F;
T 2 (outlet) t 2 (outlet)
= =
300°F 320°F
T 2 = 100 = 2.22 t1 45
tl 45 = - = 0.36 tl 125
From the chart on page 219, it is seen that one shell pass is insufficient since F is close to O. With two shell pa scs F ~ 0.90, and this arrangement would appear to be satisfactory.7 The corrected log mean temperature difference is: 0.90(L.M.T.D.) = 0.90 X 47.3 = 42.6°F The solution of this sample is also illustrated on the chart on page 218. 6 Nagle, Ind. Elly. Chern. 25, 604 (1933).
While other faclors may enter into the number of shell passes selected for a given design, allY arrangement which results in a correction factor of less than 0.80 should be rejected. 7
HEAT LOSS BY RADIATION
3.8
3.6
3.4
3.2
3.0
COEFFICIENT EMISSIVITY
Of
MATERIAL IRON OR STEEL BRIGHT OXIDIZED HIGHLY OXIDIZED COPPER POLISHED OXIDIZED BRASS BRIGHT DULL
.20-.35 .60~70
.90-.95 .10 .70 .07-.10 .25
ZINC BRIGHT .10 DULL .20 ALUMINUM PAINT .50 NON-METALLIC SURFACES BRICK,WOCO,CLOTH PAINT .95
a
2.6
1.8 THE VALUES OBTAINED FROM THESE CURVES ARE FOR IDEAL BLACK BODIES AND FOR OTHER MATERIALS MUST BE
1.6
EMISSIVITY.
1.4 1.2
1.0
.8
100
200
300
400
209
500
600
=
HEAT LOSS TO THE ATMOSPHERE BY NATURAL CONVECTION ,1 •
. I .. 1
1 ,
j
.T
t
t I
" , 1
.,
I
1.6
I
1.5
1.5 l4
=
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VALUES FROM THE CHART ARE TO BE MULTIPLIED BY THE FOllOWING SHAPE FACTORS: 1.3 VERTICAL PIPES 1.3 HORIZONTAL PIPES 1.3 VERT ICAl PLATES HORIZONTAL PLATES FACING UPWARDS 2.0 FACING DOWNWARDS 1.2
M4.
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soc. MECH. ENG. 51, 287 (1929) 240·241. MIC.GmR~~W~'~H~ILiLIIlO~onKdCjOi'tt(~'rn9~42~)~11 300
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•
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tiI_
¢ (OZG)
, . . . ,:\_
"WI!' 'oJ "lll
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300
-
I'
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r
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0-
.:
SPECiFIC HEAT OF FLUID AT AV, FLUID TEMP.-BT.U./LB.I,¥
~ <::. i _::, :::: l:~: 0;:: ~ I+~ . t Z =ABSOLUTE VISCOSITY-CENTIP()SES AT AV. TEMP. OF FLUID 20 ·ll'f:.:.;.; 4;:, ;, i: iii I::: ; II: r .Til! 'ZW'ABSOLUTE VISCOSITY-CENTIPOlSES AT AV. TUBE WALL TEMP. , iil"/ "~'I'ili!iril'I:Il;!'! II!: ".iUI HI! .l'i1I:!'I!1f:l:-:,:I::;:I:'.;.III:I':"-O:::J.,.rmillUCllI=II:liI:Jlilll::l-T l+H:tt'ti'rtJ-I' ': ,I "I.! .;. REFERENCe: S I EDER AND TATE, I NO, ENG, CHEM, 28, 1429 (1936) /0
-
I
ilJI, iili 2
[ I!;j Iii: 1111111 3
4
5 6
8
fiiUH1il~!Jl~il!H]liWltnrilllfH-'-lmrllllllif.lrrm
10
20
30
211
40 50 60 80 100
200
100 80 60 50 40
G,MASS VELOCITY -LB,/SEC /SO.FT.
C'
200
-
I' 0' INSIDE TUBE DIAMETER· INCHES
r
600 500 400
f ; 'rlt" h = FILM COEFFICIENT - a.l.U. /HR ./SQ. FT./ F. I :)-r:-/" K' THERMAL CONDUCTIVITY - B.T.U.!HR/SO. FT.!(OF PER FT.l :. ~ , ' , L :HEATED TUBE LENGTH ·fEET
. II"
j.:
-t-
4
04'
iIll .
--
-
I
-Ttl!
Fcl!-i'IIi+!-1:¥.'-P'-;.'f,lfH1!j-'
-
300400
30 20
GOO 800 1000
10
• 10
60
80 100
200
300 400
600 8001000
2000
4000 6000
8000
1000011 8000
6000 ~~# t~.:H:H. :~ ~;;1 l:-=l: ~ I' 5000 CH 1LTON AND COLBURN. IND. ENG. CHEM. 26. 1183 (L934) ...,i::rn:;- LiC 4000 .,: -j- .... t::l -':" -r- .. , ---,..--<--
-r
r
6000 5000
::t"
lEEEIEEm]4000
HEAT TRANSFER TO FLUIDS OUTSIDE TUBES
3ooo~mWi
3000
=-crlffil2000
1000 800 600 500
"'0= OUTSIDE FILM COEFFICIENT· O:r.U./HR.lSQ.FTjOF
400
K :: THERMAL CONDUCTIVITY OF FLUID atU'/SQ.FT.!HR./(OF PER FT.)
300
C ::;
00=
OUTSIDE TUBE DIAMETER - INCHES SPECIFIC HEAT OF FLUID - 8.T.U.lLB./°F
Zr::; FILM 200
#jffitl1ltm 300
VISCOSITY - CENTIPlEES
.200
{THROUGH MINIMUM FREE CROSS-SECTIONAL A~EAl - +J-'-
111,00 80
60 50 -. -i
40
MULTIPLY he BY A NeUNDLE FACTOR" AS
30
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LOGARITHMIC MEAN TEMPERATURE DIFFERENCE 100
100
90
90
80
80
80
70
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60
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55
55
50
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WHEN Ll.T, AND Ll.Tz ARE NOT WITHIN THE CHART RANGE, THESE VALUES MAY BE MULTIPLIED BY A FACTOR, AS 0.5, 2,10, OR 100. ETC. FOR EXAMPLE: GIVEN Ll. T, (ACTUAL): 200, AND Ll.TZ (ACTUAL): 20. USE 0.5 AS FACTOR, AND Ll.T,' , '00, AND .Ll.TZ: 10. FROM THE CHART, M.T. D." 39.5 OR M.T.D. (ACTUAL): 0'.5 X 39.5' 79.
REFERENCE:
217
POWER PLANT ENG. 35. 937 (1931)
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LOG MEAN TEMPERATURE DIFFERENCE CORRECTION
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t l • TEMPERATURE AT WHICH COLO FLUID ENTERS 'la II TEMPERATURE AT WHICH COLD FLUID LEAVES F
II
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Section 13
EQUILIBRIUM FLASH VAPORIZATION The vapor-liquid equilibrium relations for hydrocarbon mixtures of known analysis can be determined by trial and error from the equilibrium relations of the individual components and a material balance. For any component, i(i = 1, 2· .. , n) , (1)
Yi = Kixi
and
Xi = xiL
+ y;(100
- L)
(2)
where Yi = mole fraction of i in the equilibrium vapor Xi = mole fraction of i in the equilibrium liquid Ki = equilibrium constant of i Xi = total moles of i per 100 moles of total mixture L = moles of equilibrium liquid per 100 moles of total mixture Substituting Kixi for Yi in equation (2) and rearranging
Xi
Xi
= -L-+--'--K-i ("--l'""""OO---L-)
(3)
At equilibrium, the sum of the mole fractions in the liquid phase, x, + X2 + ... + x., must equal 1.00. While two variables, Land K" appear in the right-hand mcmber of equation (3), there are actually three variables involved sinec K, is a function of both pressure and temperature. To predict the equilibrium conditions, any two of these variables must be known and successive values oj the third assumed until the sum of the x's equals 1.00. Usually, temperature and pressure are the two variables specified, and then the trial and error involves L. Flash Vaporization of Petroleum Fractions Although the foregoing method applies to complex petroleum fractions as well as to hydrocarbon mixtures of a comparatively few known components, it has little practical significance for petroleum fractions because of the laborious calculations ...required even when component analyses are available, which is rarely the case. As a result, empirical correlations have been developed for predicting equilibrium flash vaporization curves from distillation data on crudes and petroleum fractions. The flash vaporization curve is a plot of temperature against liquid volume percent vaporized, the total vapor being in equilibrium with the unvaporized liquid at constant pressure. 222
EQUILIBRIUM FLASH VAPORIZATION /
223
:
A number of empirical correlations for determining the atmospheric flash vaporization curve have appeared in the literature, but only a relatively simple correlation would seem to be justified in view of the discrepancies between the data of various investigators. The present correllition is of the same general type as those of Piroomov and Beiswenger l and Nelson z and applies to both petroleum fractions and whole crudes. For petroleum fractions, either the 100/0 (ASTM) distillation of the fraction itself or the portion of the crude assay (True Boiling Point) distillation corresponding to the fraction may be used for predicting the flash curve. For whole crudes, the crude assay distillation should always be used in preference to the 100/0 distillation. The latter should never be used if the distillation curve flattens out below the 70'1'0 point in the neighborhood of 700°F since this is indicative of cracking. In extrapolating the atmospheric flash curves to higher or lower pressures it is suggested that the parallel method proposed by Piroomov and Beiswenger be used up to pressures of 15 psig for whole crudes and wide cuts, and up to pressures of 50 psig if the slope of the flash reference line of the fraction is not greater than 2°F/ro. By this method the atmospheric flash curve is shifted parallel to itself by a temperature interval equal to the extrapolation of the 40'1'0 point 3 on the flash reference line (FRL) as a pure compound on a vapor pressure chart. This parallel method is unsatisfactory for higher pressures, since it is known that the flash curve becomes more horizontal with increasing pressure until its slope is zero at the true critical pressure. Beyond the pressure limits recommended in the preceding paragraph for parallel extrapolation, it is suggested that a variation of the method of Watson and Nelson' be used, since no very elaborate method appears to be justified by the data. The 40'1'0 point on the FRL is extrapolated on a vapor pressure chart to a temperature 150°F above the critical temperature of the normal paraffin having the same boiling point as the 40'1'0 point. This extrapolated tempera:tur~ and corresponding vapor pressure is then used as a focal point through which straight lines are drawn on a redilinear vapor pressure chart (page 42) from the atmospheric flash temperatures for various percents vaporized. The flash curve at any pressure is determined from the temperatures at which the given pressure ordinate intersects these constant percent off (or quality) lines. These linear extrapolations do not apply if the true critical point of the fraction is approached since the copstant percent off lines become curved and converge to the true critical temperature and pressure. Piroomov and Beiswenger, Proc. API 10, No.2, Section II, 52 (1929). Nelson, "Petroleum Refinery Engineering," pp. 242-243, McGraw-Hill Book Co., New York, N.Y. (1941). 3 This i3 a slight modification of the Piroomov and Beiswenger method as they use the point of intersection between the flash and distillation curves for extrapolation. • Watson and Nelson, Ind. Eng. Chern. 26, 880 (1933). 1
2
224
DATA BOOK ON HYDROCARBONS
Reduced Crudes Perhaps the most direct method of predicting the atmospheric flash curve of a reduced crude (or .any reduced stock) which at the same time is reasonably accurate is the following: (1) Construct an atmospheric flash curve for the original crudc.
(2) Determine the number of moles of both original crude and reduced crude per given volume of original crude. (3) At the dew point (lOOro vaporized) of the original crude, assume that the reduced crude vapors are at their dew point at a partial pressure equal to their mole fraction in the total vapors (moles of reduced crude/moles of original crude) multiplied by 1 atm. (4) Extrapolate the 40% point on that portion of the flash curve corresponding to the yield of reduced crude from the partial pressure computed by (3) to 1 atm. (5) If the reduced crude has been stripped of light ends, its atmospheric flash curve is drawn through the extrapolated point parallel to the flash curve of the original crude between the abscissas corresponding to the yield of reduced crude. (6) If the reduced crude has not been stripped of light ends, a smooth curve is drawn from the split point on the flash curve of the original crude to the 20% point on the flash curve constructed by (5) to approximate the front end of the flash curve of the reduced crude. Establishment of the 20ro point as the point above which unstripped light ends cease to affect the reduced crude flash curve is, of course, entirely arbitrary but, at the same time, fairly representative. While the method outlined above is empirical to a large extent, it does have some theoretical justification. If all but one drop of reduced crude were flashed, this last drop of liquid would be in equilibrium with the reduced crude vapors at 1 atm. It is then assumed that if 100% original crude were flashed at 1 atm, the last drop of liquid would have the same composition as the last· drop of reduced crude, and the latter vapors would be at a partial pressure corresponding to their mole fraction multiplied by one atmosphere. The basis for this assumption is that the temperature difference between the boiling range of the last drop and 'that of the vapors romoved in reducing the crude is usually so great that these vapors can be considered the equivalent of steam or gas in so far as the equilibrium relations of the last drop is concerned. Making the flash curves· of reduced crudes parallel to the flash curves of their original crudes was originally suggested by Piroomov and Beiswenger and appears to be fully justified by their data. Example 1. Determine the atmospheric flash vaporization curves of an East Texas crude and its 35% bottoms (both stripped and unstripped) from the following data taken from an assay workup of t.he crude:
•
EQUILIBRIUM FLASH VAPORIZATION Assay (T.B.P.) Distillation I.B.P., of 122 5% 177 10% 262 20% 350 30% 443 40% 538 50% 636 60% 752 70% (905) 80%
Gravity °API 37.4 47.7 20.9
Crude Overhead (0-65%) Bottoms (65-100%)
Slope of DRL* =
225 Lbs/Gal 6.98 6.57 7.73
752 - 177 60 = 9.6°F/%
• DisLillaLion reference line-through 10% and 70% points.
50% Point (DRL) = 177
+ (50 -
10)9.6 = 561°F
The slope and 50% point of the flash reference line are determined from the chart on page 228: Slope (FRL) = 6.4°F/%;
50% Point (FRL) = 561 - 40 = 521°F
The atmospheric flash curve is derived from its reference line by lIsing the relation on page 229. Percent Vaporized
5 10 20 30 40 50 60
70 80
Assay Distillation (OF) Curve
DRL
t>t'
Ratio of (t>t')'s
122 177 262 350 443 538 636 752 (905) -
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.
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PREDICTION OF FLASH CURVE FROM ITS REFERENCE LINE
I
•
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II
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ASSAY (T.B.P.) DIS TlLLATION
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20
30
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40
50
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60
70
90
100
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n5
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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.
,
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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
DL50 .,I( eo U o FL K
VISCOSITY OF VAPOR - CENT'POISES DENSITY OF VAPOR - LBS.lCU. FT. 'VAPOR VELOCITY - FT.lSEC. < SIZE OF PACKING - INCHES "LIQUID RATE FACTOR <1.25 FOR BERL SADDLES
<
_1lI4.0
o 1.5
1.0
20
1
248
·.0
:0
5
.0
o
CONVERSION FACTORS TEMPERATURE To Convert From To
°e OF oR oK
OR
ox:
1.8(OC) + 459.7 of + 459.7
°C + 273.2 (OF + 459.7)1.8 °R/1.8
OF
°C
.
1.8(OC)
. (OF - 32)/1.8 . (OR - 491.7) /1.8 . oK - 273.2
+ 32
OR - 459.7 1.8("K) - 459.7
LENGTH To Convert From
To
Meters
Cm
Inches
Feet
Multiply By
1.000 100.0 2.540 30.48
Centimeters . Meters __ Inches .......................•......•.. Feet.............•......•..............
0.0100 1.000 0.0254 0.3048
0.3937 :19.37 1.000 12.00
0.03281 3.281 0.08333 1.000
AREA
To Convert From
To
Sqem
Sq em ......................•. Sqm . Sq in . Sq ft. ............•...........
1.000 10,000 6.451 929.0
VOLUME To Convert From To Cu in.
Cu ft
US gal
Sq m
Sq in.
Sq ft
M ullilJ/Y by 1.000 X10- 4 O. J550 1.000 1,550 6.451 XIO- 4 1.000 0.09290 144.0
1.076 XI0- 1 10.76 6.944 X10-1 1.000
eu em
Dbl (42's)
Imp
g&~
Liters
l>h,/liply by
Cu in Cu ft. US gal Imp gal. .. Cu em Liters Bbl (42'S). FORCE To Convert From
1.000 1,728 231.0 277.3 0.06102 61.02 9,700
5.787 X 10- 4 4.329 X 10-3 3.607 X10-3 16.39 0.01639 1.000 7.481 6.232 2.832XI0 4 28.32 0.1337 1.000 0.8326 3,785 3.785 0.160.~ 1.200 1.000 4,543 4.54a 3.531 XIO-' 2.642 X 10- 4 2.201 XIO- 4 1.000 1.000 X 10-3 0.03531. 0.2642 0.2201 1,000 1.000 5.614 42.00 34.97 1.590XIO' 159.0
To
Poundals
Pounds
1.031 XIO-4 0.1781' 0.02381 0.02857 6.290XIO-& 6.290 X 10- 1 1.000
Dynes
Grams
I a,830 4.448 X 10' 1.000 980.7
14.10 453.6 1.020XltJ" 1.000
Multiply by
Poundals Pounds Dynes Grams
. . . .
1.000 32.17 7.233XltJ' 0.07093
0.03108 1.000 2.248 X10-' 2.205 X 10-3
249
250
DATA BOOK ON HYDROCARBONS
DENSITY To Convert From
To
Sp gr Lb/gal. Lb/eu ft
. . .
Sp gr
Lo/gal
Lb/eu ft
1.000 0.1108 0.01602
Multiply by 8.3'17 1.000 0.1337
62.43 7.481 1.000
PRESSURE To Convert From
To
Lb/sq in... Lb/sq ft ... Atm ....... Kg/sq em .. In. of Hg .. Mm of Hg Ft of H,O (60°F)
Lb/sq in.
Lb/sq ft
1.000 6.944 X10-3 14.70 14.22 0.4912 0.01934 0.4335
144.0 1.000 2,116 2,048 70.73 2.785 62.43
Kg/sq em
AIm
In. of Hg
Multiply by 0.06804 0.07031 2.036 4.726XIo-' 4.882 XIO-' 0.01414 1.000 1.033 29.92 0.9678 1.000 28.96 0.03342 0.03453 1.000 1.316 XlO- 3 1.360 X10- 3 0.03937 0.02950 0.03048 0.8826
Ftof H,O (60°F)
Mmof Hg
2.307 0.01602 33.90 32.81 1.133 0.04461 1.000
51.70 0.3592 760.0 735.5 25.40 1.000 22.41
RATE OF FLOW To Convert From To
Liters per sec
Liters/sec 1.000 Gal/min. 0.06308 Gal/hr .. 1.052XI0-3 Cuft/see 28.30 Cuft/min 0.4717 Cu ft/hr. 7.862 X 10- 3 Bbl/hr .. 0.04415 Bbl/day. 1.840XlO- 3
Gal per min
Gal per hr
15.85 1.000 0.Ol667 448.9 7.481 0.1246 0.6997 0.02917
951.2 60.00 1.000 2.693 XlO' 448.9 7.481 42.00 1.750
ENERGY. HEAT. AND WORK To Convert BTU To From BTU ........... Gm-eal ......... Ft-lb ........... Hp-hr.......... Kw-hr .........
1.000 3.968XlO- 3 1.286 X 10-3 2,547 3,415
Cu ft
Cu ft
per sec
per min
Cu ft perhr
M,diiply by 0.03532 2.110 127.1 2.228XIO- 3 0.1337 8.010 3.713XIO-' 2.228 X 10-3 0.1337 1.000 60.00 3,600 0.01667 1.000 60.00 2.778 X 10-' 0.01667 1.000 1.560 X 10- 3 0.09359 5.615 6.498XlO-' 3.899XIO- 3 0.2340
Bbl per hr
Bbl per day
22.66 1.429 0.02382 641.1 10.69 0.1781 1.000 0.04167
543.8 34.30 0.5716 1.538 XIO' 256.5 4.272 24.00 1.000
Gm-cal
Ft-Ib
Hp-hr
Kw-hr
252.0 1.000 0.3241 6.417XIO' 8.605XIO'
Multiply by 777.5 3.086 1.000 1.980 XI0' 2.655XIO'
3.928 X10-' 1.558 X 10-' 5.050 X 10-' 1.000 1.341
2.928 Xio-' 1.162 Xio-' 3.767 Xio-' 0.7457 1.000
CONVERSION FACTORS
251
POWER To Convert From
To
BTU/hr .. Ft-Ib/min Ft-Ib/sec Hp...... Kw. . . . .. Kg-cal/sec G-cal/sec Tons of refrig
BTU per hr 1.000 0.07715 4.630 2,547 3,415 1.428X10' 14.28 1.2ooX10'
Ft-Ib
Ft-Ib
per min
per sec
12.96 0.2160 1.000 0.01667 60.00 1.000 33,000 550.0 44,250 737.6 1.851 X10 5 3,086 185.1 3.086 1.555X10 5 2,592
Hp
Kw
Multiply by 3.928X1O-' 2.92SX10-' 3.033X1O- 5 2.260X1O-' 1.820X1o-' 1.356X10- 3 1.000 0.7457 1.341 1.000 5.610 4.183 5.610 X10-3 4.183 X 10-3 4.712 3.514
Kg-cal
G-cal
per sec
per sec
6.999X1O- 5 5.402X10-· 3.241 X 10-' 0.1782 0.2390 1.000 0.0010 0.8400
Tons of refrig
0.06999 8.333X1o-' 5.402X1O-' 6.431X1O-' 0.3241 3.858X1O-' 178.2 0.2122 239.0 0.2845 1,000 1.191 1.000 1.191 X1o-" 840.0 1.000
INDEX Acetylenes, physical constants of, 4 Activity cOF!'f5cien~J 48 for light h)-d:ocubons b absorber oils, 67 Adiabatic comp;'0mion of gases, 82-87 Air, enthalp:i of, 182-183 specific }'ea.t cf, 88 t,hermal conductivity of, 216 viscosity of, 176 Alcohols, physical constants of, 6 Aldehydes, physical constants of, 7 Amagat's l,fr", 136-137 Area, conversion table for, 249 Aromati~s (see al~o individual compounds) physicai constants of, 5 specific gravity of saturated liquids, 142 vapor pressure of Cs, 38 viscosity of liquid, 162 A,S.T.M. distillation of petrolcum fractions, 11 average boiling points frOID, 15 equilibrium flash vaporir.ation curve from, 223, 228·-229 Avcrage boiling point3 of petrolcum fractions, 10-15 from crude ass"y (T,B.P,) distilhtions, 11 from 10% (or A.S.T.M.) distillations, i5 Benzene, enthalpy of, 112 I latent heat of vaoorization of, 77 physical constants of, 5 specific gravity of the satnrated liql\id.
Butadiene-1,3, relative volatility of, 65 spccific gmvity of the saturated liquid, 141 v::.por pressure of, 36 Butane, enthalpy of, 101 fugacity function of, 55 latent heat of vaporization of, 94-95 Mollier diagram for, 135 physical constants of, 2 relative volatility of C. hydrocarbons to, 65-66 specific gravity of the saturated liquid, 140 specific heat of vapor, 89 vapor pressure of, 30 viscosity of, 161 Butene-I, enthalpy of, 110 physical constants of, 3 relative volatility of, 65 specific gravity of the saturated liquid, 141 specific heat of vapor, 89 vapor pressurc of, 30 Butene-2, cis- and trans-, enthalpy of, 111 physical constants of, 3 relative volatility of, 65 specific gravity of the satur:l.ted liquid, 141 specific heat of vapor, 89 vapor pressure of, 30
Capacity of packed towers, 247 Carbon dioxide, enthalpy of, 182-183 physical constants of, 9 vapor ?ressure of, 37 specific heat of, 88 viscosity of, Ift2 ·~b.ermn.l conductivity of, 216 Berl saddles. 246-248 viscosity of, 176 Blending index, viscosity, 156, 173 Cnrbon monoxide, enthalpy of, 182-183 Boiling point, of hydrocarbons, 2-5 physical constants of, 9 of miscellaneous gases, 9 specific heat of, 88 of miscellaneous organic compounds, 6-7 thermal conductivity oi, 216 of petroieum fractiallS, cubic average, 11 viscosity of, 176 menu avern.ge, 10, 1<1-15 Chn.racterization factor, definition, 12 molal average, 10, 14-15 from gravity and boiling point, 16 proper average for correlating physical of typical crude fractions, 12, 17 data, 10 Columns (see Fractionating towers) volume average, 10-11 Combustion (see also Flue gas) weight average, 10, 14-15 heat of, fuel oils, 178, 180 Bubble-cap towers (see also Fractionating hydrocarbons, 2-5 towers) miscellaneous gases, 9 overall plate efficiency, 233, 245 miscellaneous organic compounds, 0-7 Butadiene-1,3, physical constants of, 3 253 1·12
254
INDEX
Combustion, heat of, paraffin and olefin gases, 178, 181 petroleum fractions, 178, 180 refinery gases, 178-179 heat available from, fuel oils, 186-188 refincry gases, 184-185 Compressibility, of gases (see P-V-T relations) of liquid petroleum fractions, 136, 143147 Compression, adiabatic, 82-87 Conductivity, thermal (see Thermal conductivity) Constants, physical (see Physical constants) Contraction, friction loss in pipes uue to, 204 Convection, heat 10SR by natural. 210 Conversion, of °A.1'.I. to specific gravity and pounds per I';allon, 138-139 of °Engler to kinematic viscosi ty, 159 of Redwood seconus to kinematic viscosity, 15X of Saybolt Furol seconus to kineml\tic viscosity, 15~-159 of Saybolt Thermo viscosity to kinelIultie viscosity, J 60 of Saybolt Lniversal seconds to kinematic viscosity, 15X tables [or, area, 2'19 density, 250 energy, heat, anel work, 250 force (weight), 249 lenl\th, 249 power, 251 pressure, 250 rate of flow, 250 temperatlll'e, 249 volume, 249 Critical pressul'C, of hydrocarbons, 2-5, 74 of miscellaneous gases, 9 of miscellaneous organic compounds, 6-7 of normal paraffins, 71 pseudo-, of light hydrocarbon mixtllres, 71 of petroleum frael-ions, 73 true, of pet.roleulll fractions, 74 Critical telllpemture, of hydrocarbons. 2-5, 69-70 of light hydrocarbons, 70 of miscellaneous gases, 9 of mis('ellaneous organic compounds, 6-7 of petroleum fractions, 72 Crude assay distillation, definition, 1, average boiling points of petroleum fractions from, 14 equilibrium flash vaporization curve from, 223-229
Crude fractions, classification of various, 13 typical, characteriz:Ltion factor of, 12, 17 gravity, °A.P.I., 1:· molecular weight of, 22-23 viscosity index of lube fractions of, 12 Cubic average boiling of petroleum fractions, 11 Cyclohexane, physical constants of, 5 vapor pressure of, 39 Cycloparaffins (see also individu:Ll compounds) physical constants of, 5 vapor pressure of, 39 Cyclopent:Lne, physical constants of, 5 vapor preSSlll'e of, 39 Dalton's L:Lw, 45, 136 Density (s('e also Specific gravity) conversion table for, 250 nitieal, hydrocarbons, 2-5 miscellaneous gases, 9 miscell:Lneous oq;anic compounds, 6-7 Dimethylacetylene, physical const:Lnts of, 4 vapor pressure of, 36 Diolefins (-,ee also individual compounds) physical constants of, 3-4 specific 1\1'lwity of satumted liquids, 141 Distillation (see A.S.T.M., Crude assay, lind Tme boiling point distillations) Efficiency of bubble-cap towers, 233, 245 Emissivity, radiant heat coefficients of, 209 Energy, conversion table for, 250 En~lel', degrees, conversion to kinematic visco!$ity, 159 Elliargement, friction loss in p;pes due to, 201 Enthalpy of, :IiI', 182-183 benzene, I 12 butane, 101 butene-I, 110 butene-2, cis- and tl'ans-, III ethane, 99 ethylene, 10~ flue gas components, CO" CO, T, etc., 182-183 heptane, 104 hexane, 103 hydrocarbon vapors, eh:Lnge with pressure, 92 isobutane, 106 isobutene, 110 isopentane, 107 methane, 98 mixtures of light hydrocarbons, 78 pentane, 102 petroleum fractions, 80-82, 114-127
INDEX Enthalpy of, propane, 100 propylene, 109 toluene, 113 Entropy (see ;\10Ilier diagrams) Equilibrium flash vaporization, of known mixtures, 222 of pctroleum fractions, 222-229 Ethanc, cnthalpy of, 99 fugacity function of, 51 latent heat of vaporization of, 94-95 MollieI' diagram for, 131 physical constants of, 2 specific gravity of the saturated liquid, 140 specific heat of vapor, 89 vapor pressurc of, 28 Ethers, physical constants of, 7 Ethylacetylcne, physical constants of, 4 vapor pressure of, 36 Ethylene, cnthalpy of, 108 fugacity function of, 50 latcnt hcat of vaporization of, 94-95 Mollicr diagram for, 130 physical constants of, 3 specific gravity of the saturated liquid, 141 specific heat of vapor, 89 vapor prcssure of, 28 Feed pbtc, optimum, fractionating towers, 234 Fenske cquation, mllumum theoretical steps at total reflux, 230 Fittings, equivalent lengths of, 193-194, 202-203 Flash vaporization, equilibrium, of known mixtures, 222 . of petrolcum fractions, 222-229 Flow of fluids, across wcirs, discharge characteristics, 205 friction factor for, 193, 198 friction loss, contraction and enlargement, 204 pressurc drop across tubc banks, 206 streamlinc, prcssure drop in pipes, 198 turbulent, equivalent Icngths of fittings, 202-203 friction factor for, 193, 198 pressurc drop in pipes, 193, 198-201 Flow of hcat (sec Heat tmnsfer) Flue gas, components, enthalpy of, 182183 percent CO, in, 189 pounds per pound of fuel, 190 thcrmal conductivity of, 192 viscosity of, 191 Force, convcrsion table for, 249
255
Fractionating towcrs (see also Fractionation) bubble cap, overall efficiency of, 233, 245 optimum feed plate, 234 packed, capacity of, 247 H.E.T.P., 2·16 prcssurc drop in, 248 Fractionation, minimum reflux ratio, 231-
233 minimum theoretical steps (Fenske equation), 230-231 theoretical steps and reflux ratio, correlation of, 244 Francis formula for rectangular weirs, 205 Friction factor, for flow of fluids in pipes, 193, 198 Fuel oils, heat available from combustion of, 186-188 heat of combustion of, 178, 180 Fugacity, of hydrocltJ'bon vapors, 62-63 of light hydrocarbons in absorber oils, activity cocfficient, 67 function of, butane, 55 ethanc, 51 ethylcnc, 50 heptaue, 59 hexane, 58 hydrogcn, 61 isobutanc, 54 isopcntanc, 56 mcthane, 49 octane, 60 pentane, 57 propane, 53 propylene, 52 Gas(es) (see also Fine gas, Refinery gas, and individual compounds) miscellancous, enthalpy of, 182-183 physical constants of, 9 spccific hcat of, 88 thermal conductivity of, 216 viscosi ty of, 176 Gasolines, vapor prcssure of, 44 Glycols, physical constants of, 6-7 Gravity, convcrsion from °A.P.I. to specific gravity and pounds per gallon, 138139 of typical crude fractions, 18 specific (sec Specific gravity) Heat, available from combustion (see Com. bustion) capacity (see Spccific heat) content (see Enthalpy) ., latent (see Latent heat of vap0rlZatlon)
256
INDEX
Heat, loss, by no.tural convcction, 210 by radiation, 209 transfer, to fluids insidc tubes, 211 to fluids outside tubes, 212 Height equivalent to tI theoret,ir.otl phte, p:teked towers, 24(J Heptane, enthalpy of, 10·1 fugacity function of, 59 latent hetlt of vaporization of, 94-95 physical constants of, 2 specific gravity of the saturated liquid, 140 specific heat of vapor, 89 vapor pressure of, 33 viscosity of, 161 Hydroe:lrbon(s) (see also individual compounds an,l Arom"tics, Olefins, etc.) critic"l temperature of, 69 light, eritic"l temperature of, 70 latent heat of vaporization of, 94-95 liquids, specific heat of, 93 physical constants of, 2-5 vapors, chauge in enthalpy with pressure, 92 fugacity of, 62-63 P-Y-T relations of, 136-137, 148154 specific hcat of, 89, 91 vapor IJre8sure of, 40 -·12 Hydrogen, fugacity function of, 61 physical constants of, 9 specific heat of, 88 thermal conductivity of, 216 viscosi ty of, 176 Isobutane, enthoJpy of, 106 fugacity function of, 54 latent heat of vaporization of, 94-95 physical constants of, 2 relative volatility of, 66 specific gravity of the saturated liquid, 140 vapor pressure of, 30 Isobutene, cnth:llpy of, 110 physical constants of, 3 relative volatility of, 65 specific gravity of the satur:lted liquid, 141 specific heat of vapor, 89 vapor pressure of, 30 Isoparaffins (see also individual compounds) molecular weight of, 20 physical constants of, 2-3 Isopenlane, enthalpy of, 107 fugacity function of, 56 lat~nt heat of vaporization of, 94-95 pbysical constants of, 2
•
Isopen~ane,
rel'!'tive volatility of, 66 speCific gravIty of the saturated liquid 140 ' vapor pressure of, 31
Ketones, physical constants of, 7 Kinematic viscosity, blending index Irom,
173 conversioll to, 15.1-156, 158--160 definition of, 155 temperature charts, 166-167 Latent heat of vtlporization, 76-77 of hydrocarbons, 76-77 of low boiling hydrocarbons, 94-95 of miscellaneous organic compounds, 6-7 of p"raffin hydroc:lrbons, 96-97 of petroleum fractions, 76-77 Length, conversion t"ble fur, 249 Log:lrithmie mean temperature diO'erence, 208, 217 correction fotctors for lllulti-pas.~ exch:lngers, 208, 218-221 Mean average boiling point, of pet,roleum fractions, 10-11, 14-15 Melting point, of hydroc:lrbons, 2-;; of inisr.ell,meous gases, 9 of miscelbncous organic compounds, 6-7 Methane, enth:llpy of, 98 fugacity function of, 49 latent he:lt of v:lporiz:ltion or, U1-95 MollieI' diagrams for, 128-129 physic:ll constants of, 2 specific gravity of thc satnrated liquid, 140 specific heat of vapor, 89 V:lpor pressure of, 27 Methy1:lcetylene, physical coustllnts of, 4 v:lpor pressure of, 35 Methylcyclopentane, physical cOIl.;tants of, 5 vapor pressure of, 39 Minimum, reflux ratio, 231.-233 theoretical fmctionating steps, 230-231 Molal avemge boiling point, of petroleum fractions, 10-11, 14-15 Molecular weight (see also Physical constants) of p:lraffins, 20 of petroleum fractions, 21 of typical crude fractions, 22-23 Mollicr diagram (s) for, butane, 135 ethane, 131 ethylene, 130 methane, 128-129
INDEX Momer diallram(s) for, propane, 133-134 propylene, 132, 134 Natural convection, heat loss to atmosphere by, 210 Nitrogen, enthalpy of, 182-183 physical constants of, 9 specific heat of, 88 thcrmal conductivity of, 216 viscosity of, 176 Octane, enthalpy of, 105 fugacity function of, 60 latent hcnt of vaporization of, 94-95 physical constants of, 2 specific gravity of the saturated liquid, 140 specific hcat of vapor, 89 vapor pressure of, 34 viscosity of, 161 Oil(s) (sec also Crude fractions and Petroleum fractions) fuel, heat "vailable from combustion of, 186-1 8 hent of combustion of, 180 lube, viscosity index of, 156, 168-172 Olefins (sec also Hydrocarbons and individunl compounds) critical temperature of, 69 hent of combustion of, 181 physical constants of, 3 specific gravity of, 141 Olefins-acetylenes, physical constants of, 4-5 Optimum fccd point, fractionating towers, 234 ' Organic compounds, miscellaneous, physical constants of, 6-7 Oxygen, enthalpy of, 182-183 physical contitants of, 9 specific hent of, 88 thermal conductivity of, 216 viscosity of, 176 Packed towers (see Fractionating towers, packed) Paraffins (sec also Hydrocarbons and individual compouuds) critical tcmperature of, 69 heat of combustiou of, lSI molecular weight of, 20 normal, critical pressure of, 71 latent heat of vaporization of, 96-97 viscosity of, 16! physical constants of, 2-3 specific gravity of, 140 Pentane, enthalpy of, 102
257
Pentane, fugacity function of, 57 latent heat of vaporizntion of, 94-95 phYSICal constants of, 2 specific gravity of the snturated liquid, 140 specific heat of vapor, 89 vapor pressure of, 31 viscosity of, 161 Petroleum fractions (sce also Crnde fractions and Hydrocarbons) average boiling points of, 10-11, 14-15 critical temperature of, 72 enthalpy of, 80-82, 11'1-127 equilibrium 11,,"h vaporizntion of, 222-229 hent of combustion of, 178, 180 Intent heat of vaporization of, 76-77 liquid, thermal conductivity of, 213 thermal ~xp:lllsion of, 136, 143-147 pseudo-critical pressure of, 73 pseudo-critical temperntuI'C nf, 72 viscosit.y-temperature charts for, 166-167 Physicnl conRt'lnts of (sec also individual compounds) acetylenes, ,1 alcohols, 6 aillehydes, 7 aromatics, 5 cycloparaffins, 5 diolefins, 3-4 ethers, 7 glycols, 6-7 isoparaffins, 2-3 ketones, 7 normal paraflins, 2 olefins, 3 olcfins-ncetylcucs, 4-5 Pipe, steel, dimCllBions of, 202 Plate efficiency of buhble-cap towers, 233, 245 Power, conversion tahle for, 251 Pressure, conversion tahle for, 250 critieat ('Pc Critic,d pres"ul'e) drop, across tube hanks, 206 due to fittings, 202 for streamline flow in pipes, 198 for turbulent flow in pipes, 198-201 in commercial pipr.s, 193, 199-201 effect of, on enthalpy of hydrocarbon vapors, 92 on viscosity of gases, 177 vapor (sec "apor.pressure) Pl'Opadiene, physical constants of, 3 specific gravity of the saturated liquid, 141 vapor pressure of, 35 Propane, enthalpy of, 100 fugacity function of, 53
258
INDEX
Propane, latent heat of vaporization of, 9495 MollieI' diagrams for, 133-134 physical constants of, 2 specific gravity of the saturated liquid, 140 specific heat of vapor, 89 vapor pressure of, 29 viscosity of, 161 Propylene, enthalpy of, 109 fugacity function of, 52 latent heat of vaporization of, 94-95 Mollier diagrams for, 132, 134 physical constants of, 3 relative volatility of, 64 specific gravity of the satumted liquid, 141 specific heat of vapor, 89 vapor pressure of, 29 Pseudo-critical pressure, 68 of mixtures of light hydrocarbons, 71 of petroleum fractions, 73 Pseudo-critical tempemture, 68 of mixtures of light hydrocarbons, 70 of petroleum fractions, 72 P-V-T rel:,tions of, hydrocarbon vapors, 136-137, 148-154 mixtures of gases, 137 R, gas constant, numerical values of, 137 Radiation, heat loss by, 209 Raoult's Law, 45 Raschig rings, 246-248 Rate of flow, conversion table for, 250 Rectification (see Fractionation) Redwood viscosity, conversion to kinematic viscosity, 158 Refinery gas, heat available from combustion of, 184-185 heat of combustion of, 178-179 Reflux ratio (scc Fractionation) Reid vapol' pressure, conversion to true vapor pressure, 4-1 Relative volatility of, C, hydrocarbons, 65-66 ethylene-ethane, 64 isopentane-pentane, 66 propylene-propane, 64 Reynold's number, e01'l'ection for equivalent lenl-(th of fittings from, 203 friction factor from, inside pipes, 198 across tube banks, 206 heat trnnsfer film coefficient from, inside tubes, 211 across tube banks, 212 Saybolt, seconds Furol, conversion to kinematic viscosity, 158-159
•
Saybolt, seconds Universal, conversion to kinematic viscosity, 158 Thermo viscosity, conversion to kinematic viscosity, 160 Specific gravity, conversion from 0 \ PI 138-139 , .. " conversion to density, 250 of aromatics, 5, 142 of diolefins, 3-4, 141 of hydrocarbons, miscellaneous gases and organic compounds, 2-9 of olefins, 3, 141 of paraffins, 2-3, 140 Specific heat of, crude fmction vapors, 90 hydrocarbon liquids, 93 hydrocarbon and petroleum fraction vapors, 91 light hydrocarbon vapors, 89 miscellaneous gases, 88 Steam, enthalpy of, 182-183 specific heat of, 88 thermal conductivity of, 216 viscosity of, 176 Steel pipe, dimensions of, 202 Streamline flow of fluids, pressure drop in pipes, 198 Temperature, conversion table for, 249 Theoretical stops, fractionating towers, 230, 233, 244 Thermal conductivity of, flue gas, 192 hydrocarbon gases, 215 liquid petroleum fractions, 213 miscellaneous gases, 216 water, 214 Thermal expansion of liquid petroleum fractions, 136, 143-147 Tolnene, enthalpy of, 113 physical constants of, 5 specific gravity of the saturated liquid, 142 vnpor pre"sure of, 37 viscosity of, 162 Towers (.,cc Fractionating towers) Tme boiling point distillation (
259
INDEX Vapor pressure or. butane, 10 butene -I, 30 butene-2, cis- nnd trans-, 30 cyclohexane, 39 cyclopentane, 39 dimethylacetylene, 36 ethane, 28 ethylacetylen~, 36 ethyl benzene, 38 ethylene, 28 gasolines, 44 heptane, 33 hexane, 32 hydrocarbons, 40-42 isobutane, 30 isobutene, 30 isopentane, 31 methane, 27 methylacetylene, 35 methylcyclopentane, 39 octane, 34 pentane, 31 propadicne, 35 propane, 29 propylene, 29 toluene, 37 vinylncetylene, 36 xylene s,38 Vaporizatiun, eqnilibrium flash (see Equilibrium flash vaporization) latent he:lt of (see Latent heat of vaporizatiun)
Vinylacetylene,
physical
consta nts
of,
4
vapor pressure or, 36 Viscosity, or aromatics, 162 of California crude fractions, 165 conversion of (see Conversion) or flue gas, 191 of gases at high pressures, 177 of hydroc arbon vapors, 174-175 of Mid-C ontine nt oils, 164 of miscellaneous gasf's, 176 of normal paraffins, 161 of Pennsy lvania crude fractions, 163 Viscosity blending index, 156, 173 Viscosity index of lube oils, 156, 168-
172
Viscosity-Temperature charts, 166-167 Volume, conversion table for, 249 Water therma l conduc tivity ot, 214 Weight, average boiling point of petroleum fractions, 10-11, 14-15 conversion table for, 249 Weirs dischal'i~e characteristics of, 205 Work; converSIOn tables for, 250 Xylenes, ]lhysi~al consta nts of, 5 specific grltvlty of the satura ted liquid, 142 vapor pressure of, 38 viscosity of, 162
