THE UNIVERSITY OF MICHIGAN INDUSTRY PROGRAM OF THE COLLEGE OF ENGINEERING VAPOR-LIQUID EQUILIBRIUM BEHAVIOR IN METHANE -HYDROCARBON -SYSTEMS Nicholas W, Prodarty A dissertation submitted in partial fulfillment of the requirements for the degree of Doctor of Philosophy in the University of Michigan Department of Chemical and Metallurgical Engineering 1967 April, 1967 IP-777

DEDICATED TO MY PARENTS

ACKNOWT.DGMENTS The author wishes to express his sincere appreciation to Professor G. Brymer Williams, under whose guidance this work was conducted, for his interest in me as a student as well as a person. To Professors D.L. Katz, J.J. Martin, J.E. Powers, and R.E. Sonntag, members of his doctoral committee, the author wishes to express appreciation for their suggestions and encouragement throughout the course of this research. I would like to express thanks to the staff of the Chemical and Metallurgical Engineering Department of the University of Michigan. Special thanks go to Mssrs. Frank Drogosz, Fanny Bolen, Doug Connell, Al Darling, Pete Severn, and John Wurster. I especially want to thank Mssrs. E.A. Daniels, AoE. Mather, and A.J. Martin for giving so generously of their time. Sincere thanks go to Miss Doris Carr and Miss Helen Walker for their constant encouragement and help in preparing this manuscript. To Donald and Jennifer Dubois I express my deepest appreciation for making life so meaningful throughout my academic career at the University of Michigan. Finally, the author would like to thank the following corporations for their most generous support: The Allied Chemical Foundation for awarding me a fellowship; The Cities Service Research and Development Company for granting me a fellowship; The Dow Corning Corporation which donated the silicone fluid; and The Phillips Petroleum Company for furnishing the pentane isomers used in this research. iii

TABLE OF CONTENTS Page ACKNOWLEDGMENT.....*OO............................... iii LIST OF TABLES.................................................. vi LIST OF FIGURES.................................................. vii NOMENCLATURE...................................... x;i ABSTRACT... a a.......... z.......... a a a a a xiv I. INTRODUCTION........ o. 0.e o 1 II. THEORETICAL CONSIDERATIONS.I..............,................ 4 III, METHODS OF PREDICTING PHASE EQUILIBRIA.....,. 9 A. Convergence Pressure......................,, 12 B. Method Using an Equation of State...........,, 12 C. Equations of State............................... 13 D. Activity Coefficients —Non-Ideal Solutions......... 19 E, Recent Developments.......................... 25 IV. MATERIALS..................... 33 V. DESCRIPTION OF EQUIPMENT.........................oo. 34 A. Hydrocarbon Loading System....................oo.. 34 B, Equilibrium System......,...........,........... 34 C, Sampling System.................................... 40 D. Composition Determination.e.........e..e...e....... 42 E. Safety..,....... e 43 VI. EXPERIMENTAL PROCEDURES.........,........... 45 VII. EXPERIMENTAL RESULTS o9* e*** **O O* O*O 51 VIII. ANALYSIS AND DISCUSSION OF RESULTS,...........oo......o.oo 89 IX. ANALYTICAL CORRELATION PROCEDURE........................... 96 A. Equation of State......................... 96 Bo Activity Coefficient......,...................... 98 C, Fugacity Coefficient of the Pure Liquid Component.. 101 X. SUMMARY AND CONCLUSIONS.ee.ge...,..e. gee..., e........... 107 iv

TABLE OF CONTENTS (CONT'D) Page REFERENCES.................................................... 109 APPENDIX A. CORRELATION OF VAPOR-LIQUID EQUILIBRIUM DATA..... 113 APPENDIX B. EXPERIMENTAL DATA................................ 135 APPENDIX C. CALIBRATIONS.................................. 147 A. Calibration of Pressure Gauge.................. 147 B. Calibration of Thermometer..................... 148 C. Calibration of Gas Chromatograph............... 148 APPENDIX D. GRAPHICAL COMPARISONS OF CALCULATED K-VALUES WITH OBSERVED K-VALUES........................... 160 V

LIST OF TABLES Table Page I Purity of Materials,,,,.. e....................... 33 II Experimental Phase Equilibrium Data for the MethaneNormal Pentane Binary System at 220~F...O........... 54 III Experimental Phase Equilibrium Data for the MethaneIsopentane Binary System at 220~F.... oo....oo o....o. 55 IV Experimental Phase Equilibrium Data for the MethaneIsopentane Binary System at 1600~F.... 0.....o.,...... 56 V Experimental Phase Equilibrium Data for the MethaneIsopentane Binary System at 280~F.,......,,.,o.,00 o 57 VI Experimental Phase Equilibrium Data for the MethaneIsopentane-Normal Pentane Ternary System at 160~F,,.... 58 VII Experimental Phase Equilibrium Data for the MethaneIsopentane-Normal Pentane Ternary System at 2200~F,.,, 59 VIII Experimental Phase Equilibrium Data for the MethaneIsopentane-Normal Pentane Ternary System at 280~F..... 59 IX Experimental Phase Equilibrium Data for the MethaneNeopentane Binary System at 1600F.0....... e o,....., 6o X Experimental Phase Equilibrium Data for the MethaneNeopentane Binary System at 220~0F.......,.,,,,..... 61 XI Experimental Phase Equilibrium Data for the MethaneNeopentane Binary System at 280~F.... o oeo.......00.o. 62 XII Experimental Phase Equilibrium Data for the MethaneNeopentane-Normal Pentane Ternary System at 160~F... e. 63 XIII Smoothed Phase Equilibrium Data for the MethaneIsopentane Binary System at 1600~F................... 64 XIV Smoothed Phase Equilibrium Data for the MethaneIsopentane Binary System at 220~FO,.. OO............... 65 XV Smoothed Phase Equilibrium Data for the MethaneIsopentane Binary System at 280~F..................., 66 XVI Smoothed Phase Equilibrium Data for the MethaneNeopentane Binary System at 160 ~F....,,.....o.O... 67 vi

LIST OF TABLES (CONT'D) Table Page XVII Smoothed Phase Equilibrium Data for the MethaneNeopentane Binary System at 220~F...000 eo..e.......o 68 XVIII Smoothed Phase Equilibrium Data for the MethaneNeopentane Binary System at 2800~F..,,,e,............ 69 XIX Graphically Determined Critical Properties for Binary Systems.... 0.0,, e a o o.e* a o. 93 XX Constants for BWR Equation of State for Individual Materials used in this Research...,.., 98 XXI Solubility Parameters.,..,................... 100 XXII Constants for Pure Components,........................ 101 XXIII Constants for Liquid Phase Fugacity Coefficient Expression................ 0 0........... a.103 XXIV MAD Computer Program for Equilibrium Vaporization Ratio Calculation,,..,,,,.,,... ooe.e........... 114 XXV Comparison of Observed and Calculated Vapor-Liquid Equilibrium Data,..,,...,,.......................... 120 XXVI Experimental Data for Binary Systems,................. 136 XXVII Experimental Data for Ternary Systems.............,, 142 XXVIII Calibration of Pressure Gauge........................ 147 XXIX Calibration of Thermometer,....... o................. 148 XXX Comparison of Analyses for Methane-n-Pentane Mixtures. 157 XXXI Comparison of Analyses for Methane-Isopentane Mix158 XXXII Comparison of Analyses for Methane-Neopentane Mixtures 0... o 0 e o o o e o.. e e. o 159 vii

LIST OF FIGURES Figure Page 1 Schematic Diagram of Experimental Equipment........... 35 2 Longitudinal Cross Section of Equilibrium Cell and Magne-Dash Stirrer..................................... 36 3 Sketch of Toepler Pump......................... 37 4 Pressure-Composition Diagram for Methane-Normal Pentane Binary System at 220~F...........,.,... e... 70 Pressure-Composition Diagram for Methane-Isopentane Binary System at 160~F..........e................ 71 6 Pressure-Composition Diagram for Methane-Isopentane Bianry System at 2200F...................6. 72 7 Pressure-Composjtion Diagram for Methane-Isopentane Binary System at 280~F.....o............. 73 8 Equilibrium Ratio-Pressure Diagram for Methane-Isopentane Binary System............................ 74 9 Equilibrium Ratio-Pressure Diagram for Methane-Isopentane-Normal Pentane Ternary System,................ 75 10 Pressure-Composition Diagram for Methane-IsopentaneNormal Pentane Ternary System at 160~F................ 76 11 Pressure-Composition Diagram for Methane-IsopentaneNormal Pentane Ternary System at 220~F...,,........... 77 12 Pressure-Composition Diagram for Methane-IsopentaneNormal Pentane Ternary System at 280~F,......e....... 78 13 Triangular Composition Diagram for Methane-IsopentaneNormal Pentane System at 160~F........................ 79 14 Triangular Composition Diagram for Methane-IsopentaneNormal Pentane System at 2200F....................... 80 15 Triangular Composition Diagram for Methane-IsopentaneNormal Pentane System at 280~F,.................. 81 16 Pressure-Composition Diagram for Methane-Neopentane Binary System at 160~F......,,..... e....... 82 viii

LIST OF FIGURES (CONT'D) Figure Page 17 Pressure-Composition Diagram for Methane-Neopentane Binary System at 220F................................ 83 18 Pressure-Composition Diagram for Methane-Neopentane Binary System at 280~F............................... 84 19 Equilibrium Ratio-Pressure Diagram for MethaneNeopentane Binary System.................... 85 20 Equilibrium Ratio-Pressure Diagram for MethaneNeopentane-Normal Pentane Ternary System...... 86 21 Pressure-Composition Diagram for Methane-NeopentaneNormal Pentane Ternary System at 160~F.......... 87 22 Triangular Composition Diagram for Methane-NeopentaneNormal Pentane Ternary System at 160~F........... 88 23 Gas Chromatograph Calibration for Methane-Normal Pentane System on a Normal Pentane Basis....... 150 24 Gas Chromatograph Calibration for Methane-Normal Pentane System on a Methane Basis.................. 151 25 Gas Chromatograph Calibration for Methane-Isopentane System on an Isopentane Basis..................... 152 26 Gas Chromatograph Calibration for Methane-Isopentane System on a Methane Basis....................... 153 27 Gas Chromatograph Calibration for Methane-Neopentane System on a Neopentane Basis........................ 154 28 Gas Chromatograph Calibration for Methane-Neopentane System on a Methane Basis.......... 155 29 Comparison of Calculated K with Observed K for MethaneIsopentane Binary at 160F........................... 161 30 Comparison of Calculated K with Observed K for MethaneIsopentane Binaryat 220~F............................ 162 31 Comparison of Calculated K with Observed K for MethaneIsopentane Binary at 280~F..................... 163 ix

LIST OF FIGURES (CONT'D) Figure Page 32 Comparison of Calculated K with Observed K for MethaneNeopentane Binary at 160F.164 33 Comparison of Calculated K with Observed K for MethaneNeopentane Binary at 2200F.0~F 1..65................ 165 34 Comparison of Calculated K with Observed K for MethaneNeopentane Binary at 2800F..................... 166 35 Comparison of Calculated K with Observed K for MethaneNeopentane-Normal Pentane Ternary at 160~............. 167 36 Comparison of Calculated K with Observed K for MethaneIsopentane-Normal Pentane Ternary at 160OF............. 168 37 Comparison of Calculated K with Observed K for MethaneIsopentane-Normal Pentane Ternary at 220~F............. 169 38 Comparison of Calculated K with Observed K for MethaneIsopentane-Normal Pentane Ternary at 280~F............. 170

NOMENCLATURE English Letters A Chromatograph peak areas A,B Parameters in van Laar and Margules equation for acitivity coefficient A~'B~ Parameters in Beattie-Bridgeman equation of state a,b,c A~,B~,C~ Parameters in Benedict-Webb-Rubin equation of state arb, c,c, Z Ao,A1,A2 A3.A4, A5 Coefficients for Chao-Seader equation for liquid phase fugacity coefficient A9 A2,B2,C2 A3, B3, C3 A A4'B4 Parameters for modified Martin-Hou equation of state A5 B5, C5 A6,B6 ab,k a,b Parameters in Van der Waals' equation of state a,b Parameters in Redlich-Kwong equation of state a,b c,d e afbgchdej Parameters in Redlich-Dunlop equation of state f,g,h,j aij Interaction parameters a Activity B Second virial coefficient C Third virial coefficient C Number of components in Gibbs phase rule c,c' Chromatograph relative response factors d Density E Energy xi

NOMENCLATURE (Cont'd) English Letters F Number of degrees of freedom in Gibbs phase rule f Fugacity G Gibbs free energy H Enthalpy H Henry's law constant K Equilibrium vaporization ratio k Boltzmann constant m Mass n Moles P Pressure P Number of phases in Gibbs phase rule Q Partition fu.nction q Effective molar volume in Wohl's equation R Universal gas constant S Entropy T Temperature V Volume x Mole fraction in liquid phase y Mole fraction in vapor phase Z Compressibility Z Effective volume fraction in Wohl's equation xii

NOMENCLATURE (Cont'd) Greek Letters (cont'd) a Coefficient of thermal expansion a Interaction parameter Y Activity coefficient 6 Solubility parameter Interaction parameter X Parameter in Wilson equation A Parameter in Wilson equation 1p Chemical potential v~ Liquid phase fugacity coefficient Volume fraction in Flory-Huggins equation Effective volume fraction Vapor phase fugacity coefficient W Acentric factor Superscripts V Refers to vapor phase L Refers to liquid phase Refers to a partial quantity, except in the case of an intensive property where it refers to a component in a mixture o Refers to vapor pressure of a pure component o Reference state E Excess property Subscripts c Refers to the critical state i,jkl Refers to enumerated species X i1 i

NOMENCLATURE (Cont'd) Subscripts (cont'd) r Refers to the reduced state T Infers total sum Refers to an extensive property on a unit mass or mole basis A Difference, final minus initial xiv

ABSTRACT The purpose of this research was to investigate the vaporliquid equilibrium behavior of methane-pentane binary systems and methane-pentane ternary systems throughout the two-phase region. Phase equilibrium was obtained in a constant-volume cell equipped with an internal stirrer. Phase compositions were determined by withdrawing small samples of each phase and analyzing them using gas chromatography. Vapor-liquid equilibrium data were obtained for the methane-isopentane binary system at pressures from about 500 psia up to the critical region at temperatures of 1600 F, 220~ F, and 280~ F. Data were obtained for the methane-neopentane binary system for pressures from about 300 psia up to the critical region at temperatures of 1600 F. 2200 F. and 280~ F. Data were obtained for the methane-normal pentane binary system at a temperature of 220~ F. The phase behavior of the methane-isopentane-normal pentane ternary system was investigated. Vapor and liquid phase compositions were determined at pressures from about 500 psia up to the critical region at temperatures of 1600 F, 2200 F, and 280~ F. Finally data were obtained for the methane-neopentane-normal pentane ternary system at pressures from 500 psia up to the critical region at a temperature of 160~ F. The experimental data obtained are discussed. Diagrams of pressure versus composition are presented for the binary systems. Diagrams of equilibrium vaporization ratio as a function of pressure xiv

are also presented. For the ternary systems, diagrams of equilibrium vaporization ratios versus pressure are presented. In addition, pressure versus composition and triangular compositions. diagrams are included. An analytical correlation which predicts equilibrium vaporization ratios is presented. Calculated equilibrium vaporization ratios are compared with observed equilibrium vaporization ratios. For the total number of points investigated in this research the absolute average deviation of the predicted equilibrium ratios from the observed ratios is within eight percent. The analytical expression, a modified form of the Chao-Seader correlation,(l0) incorporates an empirical correction factor. This factor which is based on the phase behavior of methane in isopentane effectively decreases the predicted methane equilibrium ratios by approximately 20 percent in the critical region.. The correlation represents the methane-neopentane binary system with an average absolute deviation of approximately 12 percent. Finally, a review is presented on methods employed in predicting phase equilibrium behavior. xv

I. INTRODUCTION In the absence of experimental data, reliable and accurate methods for predicting phase equilibrium behavior of multicomponent mixtures are of prime importance to the engineer concerned with the design of separation equipment, Two methods for predicting equilibrium ratios which have been widely used by the design engineer in the petroleum industry are the NGAA K-Value Charts(36) and the Kellogg Charts(27) Both methods have been used with a varying degree of success; however, each method is restricted to mixtures of paraffins, olefins, or combinations thereof. Since the development of the NGAA Charts and the Kellogg Polyco Charts, much effort has been directed toward the formulation of generalized correlations. A primary approach has been to relate the phase behavior of mixtures to experimentally determined properties of pure components and binary systems which comprise the multicomponent mixture. It would be ideal if such correlations could be related to a small number of well-behaved mathematical functions, No matter how complex, these functions could then be solved with the aid of modern, high speed digital computers. The validity of the assumptions made in deriving these functions, however, can only be ascertained by subjecting them to the test of comparison with experimental results~ A number of experimental investigations have been reported on methane and heavier hydrocarbon mixtures. Katz et al.(24) present an excellent bibliography on such systems. Equilibrium ratio data on the methane-pentane systems are incomplete. Methane, normal -1

-2pentane, and isopentane are important, naturally occurring compounds in hydrocarbon mixtures. Neopentane, because of its molecular symmetry, is of theoretical interest to the scientist whose ultimate goal is to correlate macroscopic thermodynamic functions to microscopic properties or intermolecular forces. Boomer, Johnson, and Piercey(5) have determined compositions and densities of the two-phase region at 250C and pressures ranging from 35 to 135 atmospheres for a system containing impure methane and a mixture of isopentane and normal pentane. Sage, Reamer, Olds, and Lacey(50) have experimentally determined the specific volumes of six mixtures of methane and normal pentane for seven different temperatures between 100'F and 460'F at pressures up to 5,000 psia, They have, in addition, determined the compositions of the vapor and liquid phases throughout the two-phase region for several temperatures between 100'F and 3400F and pressures from the vapor pressure of normal pentane to the critical pressure of the mixture. Experimental work on the methane-isopentane binary system has been reported by Amick, Johnson, and Dodge.(l) They report coexisting phase compositions for temperatures ranging from 1600F to 3400F and pressures from 400 psia to 1,000 psiao Their data, however, show considerable scatter, No experimental information has been found on the methane-neopentane system. Experimental work related to the pure compound neopentane (2,2 dimethyl propane) has been reported by Beattie, Douslin, and Levine(3) and more recently by Heichelheim, Kobe, Silberberg, and

-3McKetta.(18) Beattie et al. have measured the vapor pressure of neopentane from 50~C to the critical temperature, 160.60~C, and have studied the compressibility of several isotherms around the critical region in order to locate the critical point. Heichelheim et al. have investigated the compressibilities of neopentane using a standard Burnett apparatus. They have determined compressibility factors between one atmosphere and the vapor pressure at 300C to 1500C and between one atmosphere and 70 atmospheres at 161.5~C, 175~C, and 2000C. It is the purpose of this research to improve and extend existing equilibrium vaporization ratio data into the critical region on the methane-isopentane binary system and to determine the compositions of both the vapor and liquid phases throughout the two-phase region for the methane-neopentane binary system. Equilibrium ratio data for each component in the methane-isopentane-normal pentane ternary system and the methane-neopentane-normal pentane ternary system are investigated in order to extend our knowledge on more complex systems of methane in mixtures of pentane isomers. Finally, these experimental data are used to determine the reliability of a generalized correlation for predicting vapor-liquid equilibrium behavior at pressures up to the critical region.

II. THEORETICAL CONSIDERATIONS There are essentially two methods of evaluating equilibrium ratios or K values. Equilibrium ratios can be determined experimentally and theoretically. Some of the more basic experimental methods of obtaining equilibrium ratios have been reviewed by Katz et al. (24) and Sage and Reamer(49) and will not be discussed in further detail here. Recently Stalkup and Kobayashi(54a) have utilized gas liquid partition chromotography as a means of obtaining phase equlibirium data. The theoretical aspects of phase equilibria will be reviewed in this section, The difficulties encountered in attempting to obtain analytical solutions to the problem of predicting phase equilibrium behavior at high pressures will also be discussed. At the outset we give the Gibbs Phase Rule. The relationship, first stated by Gibbs, (15) is written symbolically as F+P = P+I2 (1) where F = degrees of freedom, C = number of components, and P = number of phases. The derivation of Equation (1) is given by L. 0. Case(9) and will not be repeated here~ In the present research, the number of phases is always equal to two, namely the liquid and vapor phases, Accordingly, Equation (1) reduces to: F = r- (2) In other words, Equation (2) states that the number of intensive variables needed to completely define the two-phase equilibrium system is equal to the number of components in that system

-5For a closed system at equilibrium, Gibbs has shown that the change in free energy at constant temperature and pressure is equal to zero, Expressed in equation form, d G = 0 at constant T, P. (3) We can write the free energy relation for each phase as do =v -SvT + Vvd PC + EP, /V Y- -0 (4a) AML ~LTV= + v I P +L L (4b) At constant temperature and pressure, Equations (4) reduce to Gv = Eg V Z vlt (5a),.L = L ~ad (5b) where superscripts V and L refer to the vapor and liquid phases, respectively. Since free energy is an extensive property, the total free energy (denoted by G) of the entire system is given by the sum of the free energies of the two constituent phases under consideration. We write this mathematically as: = G+vG (6) Differentiation of Equation (6) and the condition imposed by Equation (3) yields cs d GVA+ d GC=O (7)

Adding Equation (5a) and (5b) and equating the sum to zero, as in Equations (3) and (7), gives V V v V + (8) LLL (8) AL+ 5dY + 4 L/L.d +L Since the system under consideration is a closed system, the total mass of any constituent in the system remains constant, which permits writing VI + d A L =0 (9) Equations (8) and (9) imply the necessary and sufficient condition that V L (10) where the superscript bar refers to the fact that component i is in a mixture. The verbal formulation of Equation (10) is that the chemical potentials of each constituent are the same in all phases, The fugacity of a component i in a mixture is defined as: I Hi= RR dT J (SY) Integration of Equation (11) and substitution of the result into Equation (10) give the general equation =/ L 41, =, X(12) Equation (12) states the fugacity of component i in a multicomponent mixture in the vapor phase equals the fugacity of that same component in the liquid phase at constant temperature and pressure, Equation (12)

-7is then the basic equation of phase equilibrium thermodynamics. However, this equation cannot be successfully applied unless one knows the following functional relationships:'V. k* = XL[ P~ t, *@' Add at''' ] (13b) That is, for each phase one must express the functional form of the fugacity of component i in a multicomponent mixture in terms of temperature, pressure, and composition. The ultimate goal of phase equilibrium thermodynamics then is to establish relationships between thermodynamic functions, such as fugacity, and microscopic particle behavior for which the intermolecular forces are unknown quantities. This necessarily suggests the evaluation of the partition function Q, defined as: Q- I[E,/k (14) where Ei equals the energy of each of the possible states of the system. Hirschfelder, Curtiss, and Bird(22) show how the virial equation of state may be developed from the statistical thermodynamical relation between the pressure and the partition function. The range of validity of the virial equation of state is limited, however, by the convergence of the series expansion. The series expansion diverges in the liquid region. The primary value of the virial equation of state lies in the regions of low density gases and gases under moderate pressures. The problems involved in establishing a model relating macroscopic thermodynamical functions, such as fugacity, that are valid up

to the critical point of a multicomponent mixture to the molecular properties of a system are formidable, In fact, it is unlikely that statistical thermodynamics alone will be used to predict phase behavior for a long time. In the next section a review is given of the techniques, based on classical thermodynamics, utilized in calculating the fugacities of components in both vapor and liquid phases in equilibrium.

IIIo METHODS OF PREDICTING PHASE EQUILIBRIA As stated in Chapter II, the ultimate goal of phase equilibrium thermodynamics is to relate the fugacity of each component in each phase of a multicomponent mixture to the microscopic particle behavior of molecules comprising the system. In view of the level of sophistication required in the treatment of such a problem, engineers have sought solutions to this problem which require a lesser degree of sophistication and, as a natural consequence, have produced solutions which are fruitful but only approximate in light of the necessary and simplifying assumptions. A realistic goal, however, would be to relate experimental results to some mathematical functions, preferably simple ones, with a small number of constants to allow for the smoothing and interpolation of experimental data, Naturally, these mathematical functions would be based on as much of a theoretical foundation as possible to insure generality. With the aid of modern digital computers, such relations would be desirable from an engineering standpoint because attempts could be made to generalize the experimental results to such a degree that behavior of previously investigated systems, or even new systems which were not previously investigated at all, could be predicted. The content of subsequent subsections in this chapter is intended to familiarize the reader with some of the methods of predicting vapor-liquid equilibrium behavior. The advantages and limitations of the various methods presented are discussed, The pertinent equations and parameters in conjunction with the correlation used in this work are considered in Chapter IXo Some equations used in Chapter IX are first developed in this chapter. -9

-10Early attempts to predict phase behavior were dictated by the immediate needs of the petroleum industry. In the early part of the Twentieth Century, when the petroleum industry. first sought methods of predicting equilibrium ratios, the logical solution to the immediate problem was a combination of Raoult's Law and Dalton's Law, which in mathematical formulation is given simply by Equation (15). _l P_ (15) In other words, the equilibrium ratio of component i is a function of the system temperature, pressure, and component identity, but not a function of concentration. Also, Equation (15) neglects the effects of pressure on the behavior of the component in the liquid and vapor phases. It can readily be shown from classical thermodynamics for a pure component that fugacity is related to.pressure, volume, and temperature by the following relationship: eng = t -l _ P, (16) P P, Graphical integration of the right hand side of Equation (16) at constant temperature yields ratios of fugacity to pressure. Indeed, generalized plots of fugacity-pressure ratios as a function of reduced temperature and pressure have been made. Naturally these plots are as valid as the Z charts on which they are based. For mixtures, the generalized fugacities are combined with the famous Lewis and Randall Rule. (32) For an equilibrium mixture then, - = _ =,; (17)

-11Note that Equations (15) and (17) are quite similar.. The significant difference is that the fugacities partially correct for the deviations of the vapor phase from the ideal-gas law, The most important methods using the generalized fugacity concept were those of Lewis and Luke(31) and Souders, Selheimer, and Brown (54) These methods are basically the same and differ primarily by the extrapolation methods used in defining the hypothetical standard stateso Both methods are an improvement over the Dalton-Raoult Law method in that they partially correct for the pressure effect on the equilibrium ratio; however, the effect of composition is still largely neglected. In general, the equilibrium vaporization ratio Ki of component i is dependent upon pressure, temperature, and composition of both phases. The dependencies can be calculated if the pressure-volume -temperature (hereafter referred to as P-V-T) behavior is known over the entire concentration range. Because of the great shortage of P-V-T data for mixtures of interest, a great amount of effort has been expended in expressing phase behavior of mixtures in terms of pure component properties, There have been essentially two approaches to this problem. The first approach has been a purely empirical one based on available experimental results., The second approach is semi-empirical in nature, the main ingredient of which is an equation of state, The primary advantage of empirical methods is the relative simplicity and the mitigation of trial-and-error requirements. The convergence pressure technique is perhaps the most famous example of the empirical approach,

-12A. Convergence Pressure The concept of convergence pressure was perhaps first suggested by Brown et al. (54) Katz and Kurata(26) suggested the possiblity of predicting the convergence pressure of a multicomponent mixture from an equivalent binary mixture, and much of the experimental work which led to the NGAA convergence pressure charts has been done by Katz and Hachmuth. (25) The pressure at which the equilibrium ratios of each component in a multicomponent mixture appears to approach unity has come to be known as the convergence pressure. For a binary mixture, the convergence pressure is identical with the critical pressure at that temperature. Correlations using the convergence pressure concept have been published by Hadden;(l7) the Natural Gasoline Association of America, now called the Natural Gas Processors Association; Rzasa et al.; Winn; (56) Lenoir and White, (30) and Organick. (37a) B. Method Using an Equation of State The substitution of fugacity coefficient for fugacity has been found to be convenient in calculations pertaining to gaseous mixtures, The fugacity coefficient, 0i, of component i is defined by Pt~~ = >- ~~~~~(18) For gaseous mixtures the fugacity coefficient is given by the relationship (19) s tpartiss r where Zi is the partial molal compressibility factor. For equations

-13 - of state explicit in volume, one can calculate the fugacity coefficient from the following relationship: _ e A =. RT [ P7,tjJ/ ]R (20) For equations of state explicit in pressure <L [(= 1: i[r&N_ J]dv-I-z2 (21) v= V is convenient in form to calculate the fugacity coefficient i o Since no satisfactory equation of state exists for liquid mixtures, the relationship between fugacity and composition is generally expressed in terms of the activity coefficient. Hence, for a multicomponent liquid mixture, the fugacity of component i is related to pressure, temperature, and composition as: a'= rL +D~ (22) Thus, Equations (12), (18), and (22) present a theoretical basis for predicting phase equilibria, The determination and generalization of the activity and fugacity coefficients require mathematical representation in terms of parameters based on pure component properties and interaction parameters~ C, Equations of State The requirement of a good equation of state is essential in this second approach to predicting phase behavior of multicomponent mixtures at high pressure. Perhaps the most famous equation of state is that of Van der Waals, Van der Waals' two-constant equation is simply given as

-14P RT Y-L y(23) (- Z_L2 where the constant a is a measure of the cohesion between molecules and b is proportional to the volumes of the molecules. The Beattie-Bridgeman(2) five-constant equation is given as follows: p =.. RTB. -,'9 - R,~ /2 + o_.Q LA - T L R &/"7 R Ob (24) and represents experimental data with good reliability up to two-thirds the critical density. At very high pressures, above 200 atmospheres, the Beattie-Bridgeman equation fails. In the early 1940's the first real attempt was made to predict phase behavior from an equation of state. At that time, Benedict, Webb, and Rubin(4) published their equation of state. Their equation of state is a modified form of the Beattie-Bridgeman equation, The primary goal in their development was an equation to describe the phase behavior of hydrocarbon mixtures of relatively low molecular weight up to two times the critical density. The equation explicit in pressure is written in the following form: p_ Rr+ aoRTLTo-RTT2/' + -,a (25) The 3 eight constants3, A0, C a c, and y are functions of the The eight constants Bo, Ao: CoN b, a, c, a, and y are functions of the mixture composition and have been empirically evaluated by the following mixing rules: B= [ z By ] 8 ] (26a)

-15=[aen i /2)] (26b) _o (26c) -b K"[E(/ b 3s) ]3 (26da) =Ll3) X (26e) C>= [t(+sa tr ) 0 (26r) r [E(~ ffi/ r ) ] (26h) In Equations (26), the symbols with the subscript i, i.e. Boi and A refer to numberical constants for pure components, Benedict and oi, his coworkers evaluated these constants for twelve hydrocarbons~ Martin has recently modified the original Martin-HoJ35) equation. The new sixteen-constant equation describes volumetric behavior of compounds up to two and one-half times the critical density, The equation explicit in pressure is given as: _v-b (v. ~)2 - 6+,~ 6)(27) + ___ 4__ __ 74L ded3T-1- b + A-b~6 e, Equation (27) has not as yet been applied to mixtures, However, it has represented experimental data up to twice the critical density for pure components with an average deviation of 0,1 percent for Freon compounds.

Redlich and Kwong published their first equation of state in 1949i (45) The equation is of the form 2" (fish) ( if A)(28a) (28b) where i = PY6 ~Rw' (28c) R2 T 2.' (28d) b B - (28e) This equation of state contains only two constants and the authors claim satisfactory results above the critical temperature. The justification of such an equation is the degree of approximation obtained by relatively simple methods, Having acquired some degree of success, Redlich and Dunlop(44) improved upon the original equation of state by introducing a superposition function which they call the deviation function and a third parameter called the acentric factor. The equation fZ "= Z' -- Z " (29) is simply the old equation of state where Z' is Z in Equation (28a) and the superimposed deviation function is Z". The form of the deviation function is "= R [ a. (-&1 )+ (6- d7 -l )P +f2(e+f/ ) (T —/)3-/oO - iP(T-/)jT +

-17The numerical values of a, b, c, d, e, f, g, h, and j are given by Redlich and Dunlop in their article and will not be repeated here. The acentric factor is the same as Pitzer et al., (39) and is defined as Q) Xo[4' [cPO]-L oO (31) The acentric factor is determined by the vapor pressure P at the reduced temperature Tr = 0.700 and the critical pressure Pc ~ Although Equation (29) represents experimental data better than the original equation, no attempt has been made to extend the applicability of Equation (29) to the critical state or liquid region. Recently Redlich, Ackerman, Gunn, Jacobson, and Lau(43) have improved upon the Redlich-Dunlop equation and have extended its application to the liquid state and vapor pressures. The equation represents the compressibility factor as the sum of a number of terms, =ZE+,#we + L( 3 40i-Wf) (32) where Z = Root of Equation (28) u = the acentric factor L = 1 for liquids; L = 0 for gases z = Z1 (Pr Tr) Z2 = Z2 (Pr, Tr) z3 = Z3 (Pr Tr) Z4 Z 4 (Pry Tr)

-18Equation (32) has been compared to available phase equilibrium data for mixtures and pure components. The results show that hydrogen, helium, and water are not satisfactorily represented. Also the critical locus of mixtures is not well described, The virial equation of state is given by the following series expansion: P.V 1+ a(r). C(T) P T / L 2 4'(33) where the second and third virial coefficients are given by: B=CCJ e~y Be(Tr) (34a) EI L =6: 3 t+2 X( ~T ) (34b) Several potential functions are available to evaluate the second and third virial coefficients theoretically. In general, the core model of the Kihara potential(28) does a much better job of representing experimental data than the Lennard-Jones potential7(29) in particular where the molecules differ from spheres in geometry, In summary, empirical equations of state have been developed to represent experimental data to a high degree of accuracy over a limited range; that is, about twice the critical density. The virial equation of state is limited to moderate pressures well below the critical pressure simply becauseone cannot readily solve for terms higher than the third virial coefficient and the expansion series diverges on approaching the critical region, At moderate pressures, the virial equation of state can be very useful because it does give exact composition dependence,

-19The empirical equations of state are more flexible than the virial equation of state and readily lend themselves to computer programming; however, no theory exists for mixing rules. That is, all mixing rules are empirical. It appears unlikely that an empirical equation of state will be developed to predict pressure-volume-temperature and composition effects to greater than three times the critical density that can be used with any degree of facility pertaining to the evaluation of the parameters D. Activity Coefficients —Non-Ideal Solutions To account for the non-ideal behavior of liquid mixtures, it is convenient to introduce a thermodynamic function, activity~ The concept of activity has the advantage of relating non-ideal behavior from ideal behavior in one factor called the activity coefficient, The activity and activity coefficient are defined as dL (ad= Act.= g<0 (35) where fI is the reference fugacity, Many integrated forms of the Gibbs-Duhem equation at constant temperature and pressure exist, This equation at constant temperature and pressure is given as CryA~i dla,2~~~ o< r =(36) The most famous of these are, perhaps, the van Laar and Margules equations, The Carlson and Colburn(8) form of the van Laar two-constant equations for a binary mixture is given as:

-20='..... (37a) =t }E BB-otE (37b) The two-constant Margules equation for a binary mixture (given in the Carlson-Colburn modification) are of polynomial form in concentration. JS -- (24-B") 2 ~- 2 (E —4)i 2 (38a) ~Art (2B-f/)S~+ 2 (A-B)S (38b) Wohl(57) expressed the molar Gibbs excess free energy as a polynomial in liquid concentration by the expansion ATZ ~ t'z _:':- z., g,un~~~ L~ i~wS<-A~i~t uit tH(39a) where qi, qj',.0 are defined as the effective molar volumes of constituents i, j,.. and zi, zj,.. are defined as the effective volume fractions of these constituents. The term nT denotes the total number of moles in the solution. The effective volume fraction of any component i is defined as Ad = g (39b) The constants aij, aijk, aijkl are a measure of the interactions of the various components ij, ijk, ijkl comprising the liquid mixture. Expressions for activity coefficient can be obtained from the thermodynamically rigorous expression:

-21-[ i - R7J (40) The activity coefficients for a binary mixture using the threesuffix form of the Wohl equation are: t i2 [~iri L43 ( I - ) (41a) A 2= At ~[ +2 (~ Ad ) 2 (41b) Direct comparisons can be made with Wohl's equations and the previously mentioned expressions for activity coefficients. If ql/q2= 1, then it follows from the definition of z1 and z2 that Zl = xl and z2 = X2 and Rnt~= (2B8- )'62 + 2 ( )08 26 (42a) ~ -- (A2 — ) y,2 + 2 (.B-F) Z C3 (42b) Equations (42a) and (42b) are then the two-constant Margules equation. Alternately, by taking ql/q2 = a/b, then Wohl's equation reduces to the well known van Laar equationso The assumption that ql/q2 = 1 is useful in treating liquid mixtures whose constituent molecules are similar, The mathematical statement that ql/q2 = A/B may perhaps be useful in liquid mixtures with highly dissimilar moleculeso Scatchard and Hamer(53) derived equations for activity coefficients. If the effective molar volumes are replaced by the molar volumes of the pure components vL and v2, Equations (41a) and (41b) reduce to those developed by Scatchard and Hamerq Obviously the Wohl

-22equation would represent experimental data better than the three previously mentioned equations, This would be at the expense, however, of evaluating the additional constant ql/q2 The berm excess free energy was originally introduced by Scatchard(51) and is denoted symbolically as AGE. The excess Gibbs energy consists of two excess quantities —that is, an excess enthalpy and an excess entropy: ~G H IC TASE A _ A H -TAB (43) The assumption that the excess free energy is equal to zero, that is AGE = 0, leads to the concept of an ideal solution. Other less trivial assumptions would be to set either AHE or zSE = o. Most equations for activity coefficient were derived from Equation (43) using the assumption that aSE = 0 and AOHE could be written as a polynomial expansion in mole fraction or volume fraction. The condition that zSE = O leads to the concept of regular solutions.(l9) The van Laar, Hildebrand and Wood,(21) and Scatchard(51'52) equations are based on this approach. The regular solution theory of Hildebrand has been used by Chao and Seader in the development of their correlation,(10) The theory is very good for both qualitative and semi-quantitative predictions for non-polar systems such as mixtures of hydrocarbons. Scatchard made the following basic assumptions in his quantitative development of regular solutions: 1) The mutual energy of two molecules depends only on the distance between them and their relative orientation, and not at all on the nature of the other molecules between or around them or on the temperature.

-23 - 2) The distribution of the molecules in position and in orientation is random; that is, it is independent of the temperature and of the nature of the other molecules present. 3) The change of volume on mixing at constant pressure is zero, These assumptions allowed Scatchard to formulate a mathematical expression for the "cohesive energy" of a mole of liquid mixtureQ The "cohesive energy" of a binary mixture is given as: 2c,2 eCllVI9 +2eCaV, V,%aXi lyp +Ca2 VP-(44) _~we XI +_, where Cll is -E1/V1 and can be defined as the "cohesive energy density" for pure components. For multicomponent systems, Hildebrand and Scott(20) express activity coefficients in regular solutions by the following relation. RTJv rU [= \ 6] (45a) where the solubility parameter, 8, is defined as the square root of an energy density 8=[ VE.. J6 Solubility Parameter (45b) Hildebrand and Wood(2l) derived Equation (45a) by integrating the intermolecular potential energies between pairs throughout the liquid by use of continuous radial distribution functions. Equation (45a) will be discussed further in Chapter IX. The alternate approach to Equation (43) is to assume AHE = O; this leads to the concept of athermal solutionso This approach is perhaps best exemplified by Flory(14) and Huggins. (23) The Flory-Huggins equation for athermal mixtures is given by Ax_ E }anL(46a)

where xi = mole fraction and.i = volume fraction of component i o The relation between mole fraction and volume fraction is given simply as as_~~ ((46b) where Vi is the molar liquid volume of pure component i in the mixture. Recently Wilson(55) has developed a new equation to describe the variation of activity coefficient with composition. Wilson's equation is a semi-empirical extension of the Flory-Huggins equation. Wilson expresses the excess free energy at constant temperature as where A.. and A.. are adjustable parameters. Orye and Prausnitz(38) express the Wilson equation in a slightly different form. They present the Wilson equation for excess free energy as TG = E &,, < [ ZI43, ] (48a) where L )AL = 2S -L[(4AXtAt) R] (48b) and using the rigorous expression (Equation (40)) L T 7T, P, (4

-25The resulting activity coefficient for component k is: For binary systems the activity coefficients are: 14;w; AJ)$+A ia1) +A a K. (5t Oa) The Wilson equation has several advantages. First, Equations (48) present only binary constants such as Aki and Ajk, Thus, Wilson's model for multicomponent solutions requires only parameters which can be obtained from binary data which comprise the solution, Orye and Prausnitz(38) have shown the Wilson equation to give good representation of a large variety of mixtures of alcohols in non-polar solvents at low pressures. Second, the parameters Akj and Ajk have a "builtin" temperature dependence, such that one may consider (Aij - Ajj) and (\ij - \ii) to be independent of temperature over moderate temperature intervals, Eo Recent Developments A generalized correlation for the prediction of equilibrium vaporization ratios has been reported by Chao and Seader, (10) The authors claim their correlation is useful for mixtures of paraffins, olefins, naphthenes, and aromatics, Chao and Seader express the equilibrium vaporization ratio Ki in terms of rigorously defined thermodynamic functions. The expression is conveniently given as:

pi s_ A;~ (51) The method for evaluating the liquid phase fugacity coefficient v? (13) is based on the Curl and Pitzer3) modified form of the principle of corresponding states. The vapor phase fugacity coefficient is determined from the Redlich-Kwong equation of state. Finally, the liquid acitivity coefficient is based on Hildebrand's solubility parameters. C;hao and Seader state the correlation has been tested with literature data on mixtures including paraffins, olefins, aromatics, and naphthenes They state the overall deviation from 2,696 data points is 8,7 percent. The correlation has several restrictions on pressure and temperature. These are: 1) For hydrocarbons except methane -- reduced temperature: 0O5 to 1.3 based on the pure component critical temperature. pressure: up to about 2,000 psia, but not to exceed about 0.8 of the critical pressure of the system. 2) For the light components (hydrogen and methane) -- temperature. from -1000F to about 0,93 in pseudoreduced temperature of the equilibrium liquid mixture, but not to exceed 5000F. The pseudo-reduced temperature is based on the molal average of the critical temperatures of the components. pressure: up to about 8,000 psia. Grayson and Streed(l6) have extended the temperature range of the Chao-Seader generalized correlation, From new experimental vaporliquid equilibria data for high temperature (up to 8000F), high pressure

-27(3,000 psia) hydrocarbon systems, Grayson and Streed have calculated new constants for the liquid phase fugacity coefficient equation. The authors claim the new equation is useful up to 8000F for hydrogen, methane, and heavy hydrocarbons. More recently, Prausnitz, Eckert, Orye, and O'Connell(41) have presented a monograph on calculations of multicomponent vapor-liquid equilibria using computers. The vapor phase non-idealities are treated in terms of the virial equation of state truncated after the second virial coefficient, The fugacity for the vapor phase is given by Equation (19) where Xi, the vapor phase fugacity coefficient, is solved in terms of the virial equation of state. Prausnitz et al. relate the liquid phase fugacity in terms of pressure, temperature, and composition with the following equation 1A 0P_ _ (52) where yi is the pressure independent activity coefficient. Prausnitz(40) defines the reference fugacity for the light component in a multicomponent mixture by the Henry's Law constant, o so 1= C (53) For the heavy component, Prausnitz adopts the usual convention of defining the reference fugacity to be the fugacity of the pure liquid at the temperature of the solution at some specified pressure, The convention of defining the reference fugacity of the light component by Henry's Law constant offers the advantage of using a reference fugacity which can be derived from real as opposed to imaginary physical data and which is

-28not ambiguous. The disadvantage of such a convention is that it depends not only on the properties of the light component but also depends on the properties of the heavy component. In order to satisfy the Gibbs-Duhem equation at constant temperature and pressure, Prausnitz adjusts the activity coefficient [ 7 1 _ ___ (54) [ P,- T RJT such that it is a function only of composition. Using these conventions, Prausnitz defines the activity coefficient for the heavy component as (PO IJ zV dP (55a) And the pressure independent activity coefficient for the light component is defined as -~v P.- =P~) RT (55b) where (Po)'I i as 1dn1 (55c) Prausnitz states the above definitions facilitate the correlation of equilibrium data. Chueh, Muirbrook, and Prausnitz(ll) have expressed the molar excess Gibbs energy by a power series in the effective volume fraction of the solute R A(,$, +- X 2 a) - c ~... - (56a) RTQ(Y, i+- ~a ~ -- O CT 222 2..

-29where L<} FIji 3LAsp~t 92?(56b) They then determine the activity coefficients from the relations F 9 —rA& -- RrS) PT (57a) (57b) },2~ Tp P.4,~ [ 9 J'T r, p,'~, In contrast with van Laar's assumption that ql and q2 are independent of composition, Chueh et al. assume that ql and q2 are given by a quadratic function of the effective volume fraction_: OS= k~i, I[ 1 li ]2 ] (58a) v2 = YE [1 47lab2 ] (58b) where the dilation constant q.ij is a measure of how the light component swells in the liquid solution, Combining Equations (56), (57), and (58), Chueh et al. express the adjusted activity coefficients for a binary system as Rr, (,PIS) 2.AT,- ~ ~: + ~B (59a) where-7[ a] B V z (Th aV59b) =w[ /4 - + -VIC 3 / 2 where an' __

-30Chueh et al, extend their dilated van Laar model to a ternary system containing two noncondensible and one condensible components. They define the adjusted pressure independent activity coefficient for the third component as P.,,')'- H3,J f_; Jp Q d R(60) Using the molar excess Gibbs energy written as an expansion in terms of effective volume fractions 02 and O3' Rr~a1+A- a o C133 - - 2 2(M At t3 - Bee (61a) where b = $ CSZ (61b) 2 1,T, +L +X2 3 t3 and z,=- - 3 (61c) 3 let+ +X2 F y3 3 The parameter q.i is related to the liquid compositions by the relation ~ 2 + 2 ()+ 13?] (62) The mixing rules used by Chueh et al, to obtain the interaction coefficients aq123 and c23 are given as 1,23 [T(zTI3 (63a) a 23 I cX2X331] (63b)

-31Using these equations, Chueh et al. report very good agreement between calculated experimental activity coefficients for the nitrogen-oxygencarbon dioxide ternary system at 0C and predicted activity coefficients using the above mixing ruleso O'Connell and Prausnitz(37) consider the system composed of one noncondensible constituent and two condensible constituents. In an analogous manner as above, the reference fugacities for the solvents are that of the pure components, and the reference fugacity of the solute is the HenryVs Law constant in one of the solvents. By employing the unsymmetric convention for activity coefficients, they transform Wohl's method to predict the properties of a ternary system from information of the binary pairs. They describe the properties of each binary pair by a one-term Margules equation, No comparisons are made to test the validity of their derived equations, To account for the effect of pressure on the liquid phase activity coefficients, Chueh and Prausnitz(12) have just recently developed a method for predicting partial molar volumes. They first calculate molar volumes of saturated liquid mixtures from a correlation developed by Lyckman, Eckert, and Prausnitz(33) and which is based on the tables presented by Pitzero Chueh and Prausnitz give their correlation in terms of the reduced saturated volume as VQL =I) VA-(A) VL-C1)t V (64) where u is the acentric factor and V(r) V and V(2) are functions of reduced temperature and are tabulated~ Chueh and Prausnitz have fitted the tabulated values with an equation. They then calculate

-32the partial molar volumes from the relation V a [P/6 sT.- VI Yt (65) D* [c ] TiP/aV 1 and the Redlich-Kwong equation of state. The mixing rules, however, are different than those proposed by Redlich and Kwong. Comparisons between the predicted values of specific volumes of liquid mixtures and partial molar volumes and those based on experimental work are very good.

IV. MATERIALS The materials used for this study are listed in Table I. The supplier and grade of purity of each component are given, Purity analyses obtained from gas chromatograph scans are also listed. TABLE I PURITY OF MATERIALS Analysis Manufacturer s Compound of Purity Supplier Stated Purity methane 993% The Matheson Co. 99.1% normal pentane 99.9% Phillips Petroleum Co, 99,9% isopentane 99.9% Phillips Petroleum Co. 99.9% neopentane 99,8% Phillips Petroleum Co. 99.2% -33 -

V. DESCRIPTION OF EQUIPMENT The design of the equipment used in this study has been described by Brainardj (6) The equipment and modifications are described in this section. Figure 1 presents a simplified flow diagram of the apparatus used for this research. Figures 2 and 3 give a more detailed version of the equipment. The entire experimental equipment can be described conveniently in terms of subsystems. These subsystems are: the light and heavy hydrocarbon loading system, the equilibrium system, the sampling system. Also included is a discussion pertinent to the analytical technique for composition determination of the two phases, Finally, some of the safety aspects of the equipment are discussed, Ao Hydrocarbon Loading System The loading system consists of a high pressure cylinder of methane (3,500 psi) and a stainless steel micro-reaction vessel with a volume of about 140 cubic centimeters, The cylinder of methane has a high pressure regulator manufactured by the Matheson Company (Model Noo 6-670)0 The regulator is provided with 10,000 psi gauges for the inlet and discharge sides, respectively. The lines are all 1/4 inch OoDo by o0o83 inch oD,,o 316 stainless steel high pressure tubing, The valves (numbers 2,3,4 in Figure 1) are 30,000 psi items made by Autoclave Engineers, Inc. o These valves will be discussed in more detail in the following section. B, Equilibrium System Figure 2 presents a longitudinal section of the equilibrium cello The cell is a standard Aminco micro-reaction vessel (Catalogue -34

MAGNE- DASH HYDROCARBON STIRRER CHARGING CYLINDER 3000 PSI GAUGE RUPTURE / DISK / / SAMPLE BULB b TO TO TO VENT VENT VENT V M VACUUM GAS - VAPOR EXPANSION / SAMPLING -VALVE //BALL CONSTANT TEMPERATURE EQUILIBRIUM MC LEOD GAUGE NOT TO SCALE VALVE 2 IV-6 V-N VMANOMETER CYLINDER GAS CHROMATOGRAPHRECORDER-INTEGRATOR TO TOEPLER VACUUM PUMP TO VACUUM HIGH PRESSURE METHANE CONSTANT TEMPERATURE EQUILIBRIUM MC NOD GAUGET BATH CELL VALVE STOPCOCK ( THERMOMETER Figure 1. Schematic Diagram of Experimental Equipment. CLOSED-END MERCURY MANOMETER

-36RUPTURE HEAD ASSEMBLY PLUG \, —-- PRESSURE GAUGE CONNECTION UPPER SPRING MAGNETIC CORE SOLENOID HOUSING MAGNETIC CORE HOUSING \= SOLENOID SUPPORT CABLE TO TIMER LOWER SPRING. S s SSOLENOID COIL WATER OUTLET WATER INLET FEED INLET U i~ CENTERING SPRING COLLAR 3 o GASKET COLLAR NUT GLAND RETAINING RING SAMPLE OUTLETS COVER SECTION A-A BEARING PLATE VAPOR SAMPLE &\tIIl E SET SCREW OUTLET RETAINING CAP SECTION VIEW OF 1/4" AMINCO HIGH PRESSURE FITTING OPENINGS CELL BODY DASHER SHAFT UPPER DASHER DISC LOWER DASHER DISC NOT TO SCALE LIQUID SAMPLE OUTLET 1/4" HIGH PRESSURE FITTING Figure 2. Longitudinal Cross Section of Equilibrium Cell and Magne-Dash Stirrer.

-37SYSTEM TO INLET SAMPLE COLLECTION SECTION INLET f MERCURY FLOAT VALVE t INLET FLOAT VALVE ELECTRICAL CONTACT "Z" UPPER CHAMBER -9< ~~,~~ ~STOPCOCK AIR BLEEDER COMMON ELECTRICAL TO VACUUM \,-'lTOPUMP MERCURY LEVEL ELECTRICAL CONTACT Y MERCURY// STAND PIPE LOWER CHAMBER Figure 3. Sketch of Toepler Pump.

-38Noo 41-230), The cell body is made of A.I.S. Io 316 stainless steel, designed for a maximum working pressure of 11,000 psi at 1000F. The wall thickness of the cell is 5/8 inch. The cell has an approximate volume of 200 cubic centimeters and an approximate weight of 20 pounds. The cell was modified in two ways. First, the seating head, at the head gasket, was machined in such a manner that the cell would accommodate the stirring mechanism, Second, a 1/4 inch hole was drilled at the base of the cell and fitted with a 1/4 inch high pressure fitting. This fitting serves as the exit point for the liquid sample, The rate of attainment of equilibrium in the cell is increased by means of an Autoclave Magne-Dash stirrer. More specifically, agitation of the cell contents is produced by the reciprocating motion of the dasher assembly (see Figure 2). The motion of the dasher is produced by the thrust induced on a magnetic core when the coil surrounding this core is energized electrically. By using two coils, it is possible to give the dasher a positive thrust in both up and down directions, A timer which controls the flow of current to both coils (that is, energizing them alternately) controls the speed of the dasher. The duration of each stroke is then controlled by rheostats in the timer, The frequency of motion can be regulated from about one cycle per 4 seconds to 4 cycles per second, The longitudinal traverse of the dasher is approximately 1-1/2 inches, The upper and lower springs, as shown in Figure 2, act as stops for the core, Finally, the centering spring positions the dasher and supports the weight of the core, The Magne-Dash stirrer is rated for 5,000 psi operation at 650~Fo The Magne-Dash stirrer is protected by a rupture disk fabricated from 316 stainless steel and rated at 3,800 psi

-39at 72~F and 3,078 psi at 400~F. The equilibrium system, that is the equilibrium cell, is immersed in a constant-temperature bath which is equipped to maintain temperatures ranging from 100~F to 4000F. Heating is provided by six hairpin resistance heaters. Four of the heaters are rated at 1,000 watts each. One of these heaters is electrically in series with a powerstat. The fifth heater is rated at 500 watts. Constant temperatures in the bath are maintained by the sixth heater. This heater, which is rated at 300 watts, is electrically connected to a Fenwal electronic temperature indicating controller (Catalogue No. 56006). A thermistor is used as the temperature sensing probe. The probe is tied into a simple null balance bridge circuit to alternately turn the 300 watt heater off and on. The controller is provided with two modes of operation, namely on-off control operation with completely adjustable differential and proportional control with variable proportional limits. The bath fluid is silicone oil. The oil is produced by the Dow-Corning Corporation and is listed as F-i-0113 type fluid. The fluid has a viscosity of 100 centistokes at 770F. The fluid (a dimethylpolysiloxane) is usable to 500~F in open air baths. The bath is a double-walled stainless steel box with a volume of approximately 30 gallons. The oil in the bath is agitated by a mixer which is driven by an electric motor rated at 1/4 horsepower. The temperature of the constant-temperature bath surrounding the equilibrium cell and both pressure locks is determined with a calibrated mercury-in-glass thermometer (gas-filled type made by the Taylor Instrument Company, Catalogue No. 1704431). Calibration for the thermometer is given in Table XXIX of Appendix C.

The equilibrium pressure is measured with a Heise pressure gauge (Catalogue No. H-42564). The gauge is temperature compensated between -25~F to +1250F and accurate to 0.1 percent of full scale. Calibration of this gauge is given in Table XXVIII of Appendix C. C. Sampling System Vapor and liquid samples are removed from the equilibrium cell and contained by means of pressure locks. Each lock is made up of two valves (valves 5,6,7,8 in Figure 1) and a 6 inch nipple, 1/4 inch O.D. by 0.083 inch I.D., type 316 stainless steel. Sixteen gauge chromel A wire was inserted in the 6 inch nipple in order to minimize the dead volume. The valves comprising the pressure lock were a constant source of trouble due to the erosion of the stems. Originally, high temperature Autoclave valves (Catalogue No. 30VM-4071 HT) were used. However, it soon became apparent that these valves leaked after a short period of use, A new set of valves, also manufactured by Autoclave Engineers, Inc. (Catalogue No. 30VM-4071), were tried and found to be ineffectual after several openings and closings. It was finally decided that the erosion of the stems was due to "wire-drawing," A partial solution to the problem was found by using the same type valves but specifying a stellite stem as opposed to 316 stainless steel. The valve packing is glass-impregnated teflon and has proven to be satisfactory. The pressure locks are totally immersed in the constant-temperature bath fluid.

-41The vapor and liquid phase pressure locks are connected to the vapor and liquid phase expansion cylinders by high pressure 1/4 inch stainless steel tubing. The expansion cylinders used to expand the sample prior to its collection in the glass sample collecting section are gas sampling cylinders manufactured by the Hoke Corporation. The gas sampling cylinders a:re accommodated with 1/4 inch inlet fittings and 9/16 inch outlet fittings. The catalogue numbers of the vapor phase and liquid phase gas sampling cylinders are 6LD500 and 9LDl000, respectivelyo Intermediate between the expansion cylinders and the glass sample collecting section are two rupture disks rated at 107 psi at 72~F and two vacuum ball valves (see Figure 1), The ball valves are manufactured by the Jamesbury Corporation (Catalogue No, 1/2" HPV-22-GT) and rated at pressures from 0,01 microns to 4,500 psi. With the ball valves closed, the rupture disks provide a safety feature prior to admitting the sample into the glass section of the equipment if any of the high pressure valves (namely valves 6 and 7) fail. The vapor and liquid sample lines are coupled together at this point, By means of a Kovar glass seal, the all-metal system previously mentioned is connected to the glass sample collecting system, All stopcocks in the glass section are of the hollow-plug, oblique-bore vacuum type and either 4 millimeters or 8 millimeters in diameter. "Non Aq" stopcock grease distributed by the Fisher Scientific Company is used as the stopcock lubricant, Although. "Non Aq" does not possess the best vacuum lubricant properties, it is used in this research to prevent selective adsorption of the samples.

-42A Toepler pump (see Figure 3) is used to transfer the samples from the sampling lines to the collecting section. The pump is made by the Eck and Krebs Company and has a volume of about 500 cubic centimeter The glass collection section consists of expansion flasks, aground glass joint thermometer, and a closed-end mercury manometer. A cathetometer manufactured by the Central Scientific Company is used to measure mercury levels in the closed-end mercury manometer. The cathetometer is capable of discerning distances as small as 50 microns. D. Compos it ion Determinat ion A Perkin-Elmer Vapor Fractometer (Model No. 154-D) equipped with a thermal conductivity cell as the sensing device is used for the separation and analysis of the vapor and liquid samples. Essentially, a carrier gas, in this case helium, and the sample pass through a column where the sample components are separated. The column used in all aspects of this research is 14 feet of 1/4 inch tubing packed with squalane (20 percent) on a Chemisorb support. The samplercomponents are swept one by one into the sensing side of the detector. Both sides of the thermal conductivity cell are incorporated into a balanced bridge circuit. When a thermal conductivity change occurs between the reference gas and the sample plus reference gas, a resulting bridge imbalance provides a voltage which drives the pen on a recording potentiometer. The recorder is a Leeds and Northrup Model G recording potentiometer. The recorder has a one second, full scale balance time and a 5 millivolt nominal span.

-43The gas sample is introduced into the Fractometer by means of a Perkin-Elmer precision gas sampling valve (Catalogue No, 008-0659). The valve is made of stainless steel and teflon with a sample volume of about 2 cubic centimeters. A Perkin-Elmer printing integrator (Model No. 194-B) is used to integrate the area under the resultant chromatographic curves. The integrator is a standard velocity, servo-computing arrangement with the input signal produced by a potentiometer installed on the recorder shaft, Within the integrator, an amplifier drives a servo-motor coupled to a tachometer generator and to a printing counter which registers the total number of shaft turns accumulated, The tachometer generator produces an output voltage which is linearly proportional to the speed at which it is driven by the servo-motor, The amplifier compares the tachometerproduced signal to that of the potentiometer in the recorder and continuously regulates the speed of the servo-motor, Each value of the potentiometer input signal corresponds to a definite servo-motor and tachometer speed, Since the rate of rotation of the printing counter is proportional to the recorder pen position, the total number of turns registered in a given time interval is proportional to the integral of the recorder pen position during the same time interval, Hence, it is proportional to the area under the curve produced by the pen, The Perkin-Elmer integrator has a maximum integrating rate of 6,000 counts per minute, Eo Safety A conscientious attempt was made to incorporate safety features into the design of the experimental equipment. First, the

-44equilibrium cell and pressure locks were hydrostatically tested to 3,000 psi prior to any experimental runs. A safety shield fabricated of 1/4 inch steel plate surrounds the constant-temperature bath and the equilibrium cell contained therein. A panel was cut out on one of the sides of the barricade such that the valves can be manipulated with a minimum exposure of the operator to the high pressure equipment. As previously mentioned, the Magne-Dash stirrer and sampling lines are provided with rupture disks. Finally, a hood was placed over the equilibrium constanttemperature bath, such that in the case of a rupture disk failure the contents of the cell would be transported out of the room to the exterior of the building.

VI. EXPERIMENTAL PROCEDURES In the description of the experimental procedure, references to all component equipment identity are made to Figure 1, The initial startup procedure will be discussed in detail. The entire experimental equipment is evacuated to a pressure of 10 microns or less for a period of not less than 24 hours. With reference to Figure 19 at the outset all stopcocks and valves are open with the exception of valves 2,3,9, and 10o Valve 1 denotes the valve on the high pressure cylinder of methane, The high pressure line between valves 1 and 2 has been flushed several times previously with methane from the cylinder, Heavy hydrocarbon (pentanes) loading is accomplished by pipetting a prescribed amount, about 90 cubic centimeters, of pentane (normal or iso-) into the charging cylinder, Since neopentane boils at a temperature considerably below room temperature, it is charged into the equilibrium cell directly from the containing cylinder, which is placed on a pan balance, In addition, the equilibrium cell is cooled down by direct contact with solid carbon dioxide, About 50 grams of neopentane are normally administered to the cello In charging, valves 6 and 7 are closed and valve 3 is then opened to permit the pentane to enter the equilibrium cell by gravity and pressureinduced flow. Valves 3 and 4 are then closed, Valve 6 is opened and closed several times, thereby evacuating any air that is dissolved in the pentane, Closing valve 6, methane is then transferred into the equilibrium cell by setting the pressure regulator at 3,500 psi, opening valve 2, and cracking valve 4, -45

-46Upon reaching a pressure (as indicated by the Heise pressure gauge) somewhat less than the desired operating pressure, valve 4 is closed, The Magne-Dash stirrer is then initiated and operated at about 3 cycles per second, The electrical heaters, including the Fenwal temperature controller, are energized and the oil bath agitator is turned on, Having reached the desired operating temperature, the desired operating pressure is obtained by venting some vapor from the cell or pressurizing the cell with methane from the high pressure cylinder. Once the oil bath has reached the desired operating temperature, the heat input into the cell is adjusted by means of a variac and the temperature controller. At this point, the cell and its contents are allowed to physically equilibrate for no less than eight hours. It was found during preliminary runs that a minimum of about four hours was required to attain physical equilibrium, and about eight hours was required to attain equilibrium near the critical region for the cell geometry, During this equilibration time, the stirrer is operated at a frequency of about one cycle per second, The bath thermometer is checked many times to determine the constancy of the temperature indication, The first step in the vapor sampling procedure is to turn off the stirrer and record the temperature and pressure readings, Valves 5, 135 and 14; both vacuum ball valves; and stopcock "e" are then closed, Valve 6 is opened and closedo In so doing, a vapor sample is transferred from the cell to the previously evacuated pressure lock, Valve 5 is then opened and the vapor sample is allowed to expand into the vapor expansion cylinder, Ball valve 1 is opened and the vapor sample is further expanded into the sample collecting section, At this point, stopcock "d" is closed and the Toepler pump initiated,

The Toepler pump (see Figure 3) operates in such a manner that the sample is transferred from the sampling lines to the sample collecting section, In actual operation, the mercury is in the lower chamber at the outset, With stopcock "d" closed, air is admitted to the lower chamber of the Toepler pump by way of a bleed valve, The air forces the mercury contained therein through a standpipe into the upper chamber, The mercury rises in the upper chamber, forcing the gas sample through the mercury float valve "t", When the mercury makes contact with electrical contact "Z", a relay is closed which automatically starts a vacuum pump, thereby evacuating the lower chamber and draining the mercury back into it, The sample, which has been transferred to the sample collection chamber, is now contained by the one-way mercury float valve "t", As the mercury fills the lower chamber, the gas standpipe is again connected to the upper chamber of the pump. Mercury makes contact at point "Y" and the relay is opened, thereby shutting off the vacuum pump, The cycle starts over again by introducing air by way of the bleed valve into the lower chamber, Seven to ten cycles of the Toepler pump are found to be sufficient to move the gas sample from the expansion cylinder and the sampling lines into the sample collection chamber, The temperature of the sample is measured with a thermometer, and pressure measurements are performed with the closed-end mercury manometer and the aid of a cathetometer, Normally, the pressure of the sample is maintained at pressure between 10 and 15 centimeters of mercury by using an appropriately sized sample bulb "b" o During all preliminary runs the appropriate volume of sample bulb "b" was determined using the criterion of a maximum pressure of 20 centimeters of mercury. No problems of partial condensation of the samples were encountered in this research,

Analyses of both liquid and vapor samples are carried out in the vapor phase, The analytical procedures involve the introduction of the sample into the gas chromatograph by means of a gas sampling valve. About one hour before the vapor sample is withdrawn from the cell, the gas chromatograph, recorder, and integrator are turned on, The helium carrier-gas flow rate is adjusted to a flow rate of 72 milliliters per minute (at 250C, 740 millimeters of mercury), as determined by a soap film gas meter, The chromatograph oven temperature was maintained at 80~Co The sample loop is evacuated by means of a vacuum pump, Stopcock "g" is closed and stopcock "f" opened, Once the sample pressure remains constant, the gas sample is introduced into the chromatograph by turning the gas sampling valve, which switches one of the two sample valve tubing loops into the flowing carrier-gas stream, The resulting chromatographic areas are then determined by using the Perkin-Elmer printing integratoro A minimum of two analyses are made for each sample, In all cases, duplicate samples differed less than,75 percent, The sample loop is evacuated and anew sample is introduced in the same manner as previously described, During the time that the vapor sample is being transferred from the sample lines to the sample collection chamber by means of the Toepler pump, the liquid sampling line is flushed. This is done to acquire a representative liquid sample from the cell, since preliminary runs showed that the liquid in the liquid drawoff tube was not of the same composition as that in the equilibrium cell The flushing procedure is accomplished in the following manner. Valves 8, 12, and 14

are closed and valve 10 is opened, A liquid sample is introduced into the liquid pressure lock by opening and closing valve 7, Valve 8 is opened and the liquid sample vented, This procedure of opening and closing valves 7 and 8 is repeated four times, Preliminary runs revealed that the fifth sample was representative of the liquid equilibrium composition in the cello Geometrically, four flushings are equivalent to about 1,5 times the volume of the liquid drawoff line, Valve 10 is closed and valves 8 and 14 are then opened. Having analyzed the vapor phase composition, the system is prepared for analysis of the liquid phase composition, A larger sample bulb "b" is inserted into the system after closing stopcocks "c" and "d"o For vapor samples, the sample bulb volume ranges from about 25 to 250 cubic centimeters. For liquid samples, the sample bulb volume ranges from approximately 250 to 1000 cubic centimeters. The entire system is then evacuated to a pressure of about 10 microns for a period of about one-half houro Liquid samples are withdrawn in the same manner as vapor samples. The valve manipulations in withdrawing and collecting the liquid sample are analogous to those of the vapor sample, In fact, they are symmetrical from the vapor sampling line up to the point where the vapor and liquid sample lines merge, The stopcocks in the glass section are opened and closed in the manner used when analyzing the vapor sample, In collecting the liquid sample, about twice the number of cycles are required to transfer the liquid sample from the sampling lines to the sample collection chamber as compared to the pumping time for the vapor sample, In addition, the liquid sample is allowed to expand, and it is collected again to insure mixing. This mixing process is accomplished in the

-50folowing manner. After the sample is collected, stopcock "at is closed and stopcock "d" opened. The sample then expands into the upper chamber of the Toepler pump. Stopcock "d" is closed and the pumping procedure repeated until all of the sample is once again contained in the sample collection chamber, Analysis of the liquid sample is accomplished in an analogous manner to that for the vapor sample. As with the vapor phase samples, duplicate analyses are run as standard procedure. The equipment is then prepared for the next run by evacuating the entire system, Methane is administered into the equilibrium cell, thereby increasing the pressure. The Magne-Dash stirrer is reactivated and the equilibrium cell is allowed to re-equilibrate for at least eight hours, For runs near the critical region the cell is allowed to equilibrate for about twelve hours. In the course of this research, experimental data were obtained at three isotherms, namely 1600F, 2200F, and 280 oF.

VII. EXPERIMENTAL RESULTS The experimental data obtained in this research are presented in Tables II through XII and Figures 4 through 22 in the same order as the data were taken chronologically. The experimental phase equilibrium data for the methanenormal pentane system are given in Table II. Experimental data were taken at one temperature (220~F), and several points were taken at essentially the same pressures to establish the reproducibility of the entire system. The experimental data on this system are compared with those of Sage, Reamer, Olds, and Lacey(50) in Figure 4. As can be seen, agreement between Sage et al, and this work is quite good. Tables III through V present the experimental phase equilibrium data for the methane-isopentane binary system. Figures 5 through 7 show the pressure composition data at temperatures of 1600F, 220~F, and 280~F, respectively. Included in Figures 5 through 7 are the experimental data reported by Amick, Johnson, and Dodge,(l) Figure 8 is a loglog plot of the equilibrium vaporization ratios, K, of methane and isopentane as a function of pressure. All three isotherms (160oF, 220OF and 2800F) are included in Figure 8, The loci in Figure 8 represent smoothed equilibrium vaporization ratios for methane and isopentane, Smoothed K values were obtained from Figures 5 through 7 and are presented in Tables XIII through XV. The uncertainty of the smoothed values of methane composition is believed to be + 0~002 mole fraction, Experimentally determined K values for methane and isopentane are also presented in Figure 8, -51

-52Tables VI through VIII present the experimental phase equilibrium data for the methane-isopentane-normal pentane ternary system. The Gibbs Phase Rule states that the number of intensive variables required to specify the system is three for a two-phase system containing three components. The intensive properties selected to determine the ternary system are pressure, temperature, and isopentane concentration to isopentane plus normal pentane concentration in the liquid phase, Figures 9 through 15 graphically present the experimentally determined phase equilibrium behavior of the methane-isopentane-normal pentane ternary system, Figure 9 illustrates the equilibrium ratios for methane, isopentane, and normal pentane as a function of pressure for the three isothermal conditions investigated in this research, Figures 10 through 12 present the pressure-composition diagrams for the methane-isopentane-normal pentane ternary system at the three tempertatures of 1600F, 2200F, and 2800F, respectively, The loci of Figures 9 through 12 are described by a parameter of isopentane concentration to isopentane plus normal pentane concentration, Figures 13 through 15 give the ternary composition diagrams for three pressures. They show the decrease of the two-phase region with increased pressure, They also illustrate the small change of the liquid phase mole fraction parameter with pressure for different isotherms, The data for the methane-normal pentane binary system in Figures 13 through 15 are those of Sage et alo(50) Tables IX through XI give the experimental phase equilibrium data for the methane-neopentane binary system. Also tabulated are the experimentally determined K values of methane and neopentane, Figures 16 through 18 are the pressure-composition curves for the three isotherms

-53at which experimental data were obtained. Figure 19 compares smoothed values of the equilibrium vaporization ratios of methane and neopentane with the experimental values as a function of pressure, The loci of the three isotherms presented in Figure 19 represent smoothed data as determined from Figures 16 through 18, These smoothed data are presented in Tables XVI through XVIII No data have been found in the literature for the methane-neopentane binary system. Hence, no comparisons are made with this work~ Table XII presents the experimental phase equilibrium data for the ternary system of methane-neopentane-normal pentane. Experimental data were obtained for one isotherm, namely 160~F, Figure 20 is a plot of the equilibrium vaporization ratios of methane, neopentane, and normal pentane at 160'F as a function of pressure on logarithmic coordinates, Figure 21 shows the pressure-composition diagram for the methane-neopentane-normal pentane ternary system at 1600F. The locus in Figure 21 is described by a parameter of neopentane concentration to neopentane plus normal pentane concentration in the liquid phase, Finally, Figure 22 presents a ternary composition diagram of this system. It demonstrates the shrinkage of the two-phase region with increased pressure and the relative independence of the heavy components with pressure on a methanefree basis for the liquid phase. Methane-normal pentane data illustrated in Figure 22 are those of Sage et al.(50)

TABLE II EXPERIMENTAL PHASE EQUILIBRIUM DATA FOR THE METHANE-NORMAL PENTANE BINARY SYSTEM AT 220 F Vapor Phase Composition Liquid Phase Composition Equilibrium Ratio Pressure (mole fraction) (mole fraction) Run no, (psia) methane n-pentane methane n-pentane methane n-pentane 21 1502 0,808 o,192 0.380 0,620 2o13 0.310 22 1265 o,812 0o 188 0o 324 0.676 2o 50 0.279 23 1231 0,810 0o190 0o306 o.694 2.65 0.273 24 1023,806 0,194 0.253 0.747 3.18 0.260 25 1001 o0.805 0.195 0.247 0.753 3.26 0.259 26 1999 0,740 0.260 0.532 0.468 1.391 o.556 27 1777 0,788 0o212 0o 456 0.544 1.729 0.389 28 1501 o,808 0.192 0,382 0.618 2.11 0.310 29 1260 0o816 0o184 0.310 o0,690 2.63 0.267 30 1005 00,814 0,186 0,248 0.752 3.28 0.248

TABLE ilIl EXPERIhENTAL PHASE EQUILIBRIUM DATA FOR THE METI{AiNE-ISOPENTANE BINARY SYSTEM AT 220"F Vapor Phase Composition Liquid Phase Composition Equilibrium Ratio Pressure (mole fraction) (mole fraction) Run no, (psia) methane i -pentane methane i-pentane methane i-pentane 51 1256 0,788 0,212 0,331 0,669 2,58 0,317 52 1503 0,774 0.226 0.396 0,60o 1,955 0,3711 33 1721 0,7746 0.254 01.45 0.546 1.615 0,146 34 1899 o,686 0.314 0,566 0. 434 1,212 0.721 55 1001 0,791 0,209 0,262 0,758 5.02 0.281 36 759 0,765 0.235 0,192 0,808 5,99 0.290 57 199 0,710 0,290 0,118 0,882 6,01 0.528

TABLE IV EXSPERIMENTAL PHASE EQUILIBRIUM DATA FOR THE METHANE-ISOPENTANE BINARY SYSTEM AT 160 0F Vapor Phase Composition Liquid Phase Composition Equilibrium Ratio Pressure (mole fraction) (mole fraction) Run no, (psia) methane i-pentane methane i-pentane methane i-pentane 58 502 o,841 0,159 o,142 0.858 5,93 0.1856 59 755 0,872 0,128 0.218 0.782 3.99 0.16h1 40 1001 0,885 0,115 0,283 0,717 5.15 01599 41 1255 0,879 0.121 0.351 0.649 2.50 0,1859 42 1505 0,869 0.131 o.418 0,582 2.08 0,225 43A 1759 0,855 0.147 0.489 0.511 1.744 0.288 44 1992 0,821 0,179 0.545 0,455 1.506 0.593 45 2191 0,7141 0.259 o,633 0,567 1.170 0.705

TABLE V EXPERflADNTAL PHASE EQUILIBRIUM DATA FOR THE METHAUNE L-ISOPENTAINEL BINARY SYSTEM AT 2800F Vapor Phase Composition Liquid Phase Composition Equilibrium Ratio Pressure (mole fraction) (mole fraction) Run no, (psia) methane i-pentane methane i-pentane methane i-pentane 46 511 0,520 o,48o 0.092 ogo8 5.68 0,528 47 o.759 O,65 0.597 0161 0.839 3.74 o. 47y 48 1001 o,066 0,364 0.231 0.769 2,75 0.474 49 1267 0,651 0,349 0.515 0.685 2,07 0,510 49A 1277 0,643 0.357 0.330 0.670 1.948 0.555 50 1517 0,581 0.419 0,488 0,512 1.191 0.818

TABLE VI EXPERlI1ENTAL PHASE EQUILIBRIUM DATA FOR THE METHANE-ISOPENTANE-NORMAL PENTANE TERNARY SYSTEM AT i6050 Vapor Phase Composition Liquid Phase Composition Equilibrium Ratio Pressure (mole fraction) (mole fraction) Run no. (psia) methane isopentane n-pentane methane isopentane n-pentane methane isopentane n-pentane 51 504 0.871 0.040 0.089 0.138 0.223 0.639 6.29 0.1775 0.1397 52 755 0.894 0.030 0.076 0.211 0.206 0.583 4.25 0.1484 0.1296 m 53 1005 0.897 0.030 0.075 0.274 0.189 0.537 5.28 0.1567 0.1365 54 1495 0.889 0.051 0.080 o.406 0.153 o.441 2.19 0.200 o.1821 55 1975 0.849 0.040 0.111 0.504 0.129 0.367 1.685 0.314 0.302 55A 1995 0.842 0.042 o.116 0.521 0.125 0.556 1.616 0.343 0.526 58 2268 0.758 0.063 0.178 0.593 0.104 0.303 1.278 0.608 0.589

TABLE VII EKPER=EHqTAL PHASE EQUILIBRIUM DATA FOR THE METHANE-ISOPENTANE-NRIAAL PENTANE TERNARY SYSTE,~ AT 220 F Vapor Phase Composition Liquid Phase Composition Equilibrium Ratio Pressure (mole fraction) (mole fraction) Run no. (psia) methane isopentane n-pentane methane isopentane n-pentane methane i sopentane n-pentane 59 1765 0.771 0.061 o.168 0.454 0.139 0.407 1.698 o.44i 0.413 60 2047 0.747 o.065 o.i88 0.555 0.112 0.333 1.346 0.584 0.564 61 1519 0.801 0.053 o.146 0.389 0.154 0.457 2.06 0.342 0.321 62 1265 0.810 0.051 0.139 0.319 0.172 0.509 2.54 0.296 0.273 63 995 0.810 0.051 0.139 0.251 0.188 0.561 5.22 0.272 0.248 64 755 0.788 0.058 0.154 0.186 0.205 0.609 4.22 0.282 0.254 65 507 0.744 0.070 0.185 0.120 0.220 0.660 6.19 0.319 0.281 TABLE VIII EXPERIMENTAL PHASE EQUILIBRIUM DATE FOR THE METHABE-ISOPENTANE-NORMAL PENTANE TERNARY SYSTEM AT 280~F Vapor Phase Composition Liquid Phase Composition Equilibrium Ratio Pressure (mole fraction) (mole fraction) Run no. (psia) methane isopentane n-pentane methane isopentane n-pentane methane isopentane n-pentane 66 559 0.563 0.116 0.320 0.099 0.220 0.681 5.64 0.529 0.471 67 760 0.630 0.097 0.275 0.162 0.202 0.636 3.89 0.479 0.429 68 1001 0.662 0.087 0.251 0.231 0.187 0.582 2.87 0.464 0.431 69 1253 o.674 0.081 0.245 0.306 0.164 0.529 2.20 0.492 0. 463 66A 541 0.568 0.120 0.512 0.105 0.252 0.665 5.51 0.516 0.469 67A 757 0.629 0.102 0.269 0.164 0.214 0.622 3.83 0.476 0.452 68A 1051 0.665 0.091 0.244 0.242 0.195 0.563 2.75 0. 465 0. 434 69A 1255 0.674 0.087 0.238 0.504 0.180 0.516 2.22 o.484 0.462 70B 1565 0.616 0.099 0.285 0.453 0.138 0.408 1.558 0.718 0.698

TABLE IX EXPERIMENTAL PEASE EQUILIBRIUM DATA FOR THE METHANE-NEOPENTANE BINARY SYSTEM AT 1600F Vapor Phase Composition Liquid Phase Composition Equilibrium Ratio Pressure (mole fraction) (mole fraction) Run no, (psia) methane neopentane methane neopentane methane neopentane 71 511 0,761 0,239 0.153 0,847 4o96 0, 282 72 763 0o,797 0,203 0.232 0,768 3~43 0.264 73 1005 0, 819 0o.181 0.312 0.688 2.63 0,263 74 1273 0.813 0o,187 0,391 0,609 2.08 O.306 74A 1281 0,813 0,187 0,398 0o602 2.04 0,311 75A 1521 0,784 0,216 0o482 0,518 1,628 0o,416 76B 1709 0,727 0,273 o,560 0,440 1.298 0, 620 77 1748 0,685 0.315 0,o603 0.397 1.137 0.792 82 310 0.667 05333 0,085 0,915 7.83 0,364

TABLE X EXPERIMENTAL PHASE EQUILIBRIUM DATA FOR TEE METHANE-NEOPENTANE BINARY SYSTEM AT 2200F Vapor Phase Composition Liquid Phase Composition Equilibrium Ratio Pressure (mole fraction) (mole fraction) Run no, (psia) methane neopentane methane neopentane methane neopentane 83 308 0,395 0o605 0,051 0.949 7.81 0.637 84 503 0 563 0 437 0,117 0,883 4.82 0o 495 85 748 0.o 639 0 361 0o 197 0o 803 3024 0. 450 86 1008 0,o 670 0,330 0.282 0.718 2.38 0.o 459 87A 1251 0,654 o0346 0,377 0.623 1.735 0.556 88A 1434 0,585 0o 415 0o 471 0.529 1.242 0,784

TABLE XI EXPERIMENTAL PHASE EQUILIBRIUM DATA FOR THE METHANE-NEOPENTANE BINARY SYSTEM AT 2800F Vapor Phase Composition Liquid. Phase Composition Equilibrium Ratio Pressure (mole fraction) (mole fraction) Run no. (psia) methane neopentane methane neopentane methane neopentane 91 506 0,280 0.720 0.068 0.932 4.i0 0.772 92 755 09407 0.593 0.163 0.857 2949 0.709 93B 1004 0,416 0,584 0.281 0.719 1.479 o.8

TABLE XII EXPERINMENTAL PHASE EQUILIBRIUM DATA FOR THE METHANE-NEOPENTANE-NORMAL PENTANE TERNARY SYSTEM AT 160~0F Vapor Phase Composition Liquid Phase Composition Equilibrium Ratio Pressure (mole fraction) (mole fraction) Run no. (psia) methane neopentane n-pentane methane neopentane n-pentane methane neopentane n-pentane 95 503 0.845 0.058 0.097 0.141 0.216 0.643 6.01 0.266 0.1512 96 751 0.871 o.o46 o.o83 0.206 0.200 0.594 4.23 0.230 0.1391 97 1251 0.878 o.o040 0.082 0.338.164 0.498 2.60 0.246 0o.64i 98 1505 0.870 o.o,4 0.089 0.400 0.148 0.452 2.18 0.276 0.1972 99 1759 0.855 0.043 0.102 0.461 0.131 0.408 1.855 0.327 0.251 100 2013 0.805 0o.o054 0.141 0.550 0.111 0.339 1, 44 0. 482 0. 417 101 2120 0.775 0.059 0.166 0.602 o.og98 0.300 1.,S8 0.606 0.551 102 1006 0.879 0.040 0.081 0.278 0.176 0.546 3.16 0.231 0.1479

TABLE XIII SMOOTEKD PH/ASE EQUILIBRIIUM DATA FOR TIE METHANE- ISOPEWTANE BINARY SYSTEM AT 160~F Vapor Phase Composition Liquid Phase Composition Equilibrium Ratio Pressure (mole fraction) (mole fraction) (psia) methane i-pentane methane i-pentane methane i-pentane 500 0, 840 o,160 0.139 o 861 6.04 o,186 750 0 872 0,128 0 216 0,784 4,04 0.163 1000 0,884 0 o 116 0. o 285 0.715 3.10 O. 162 1250 0,o 880 0 o,120 0.350 0. 650 2. 51 0, 185 1500 0,o 874 -0126 0o 417 0583 2, 10 0,216 1750 0o 853 0.147 0,483 0.517 1.77 0.284 2000 0o,817 0o,183 0.549 0,451 1,49 0.406 2150 o 768 0,232 0.605 0o 395 1.27 0.587 2213 0o,688 0,312 0.688 0.312 1.00 1.00

TABLE XIV SMOOTHED PHASE EQUILIBRIUM DATA FOR THE METHANE-ISOPENTANE BINARY SYSTEM AT 2200F Vapor Phase Composition Liquid Phase Composition Equilibrium Ratio Pressure (mole fraction) (mole fraction) (psia) methane i-pentane methane i-pentane methane i-pentane 500 0.710 0,290 0,119 0,881 5.97 0.329 750 0,766 0,234 0,195 0,805 3.93 0.291 1000 0,790 0,210 0,262 0.758 5,02 0,284 1250 0.789 0,211 0,330 0,670 2.39 0.315 1500 0,774 0,226 0,596 0,60k 1,95 0.374 1750 0,742 0,258 o,465 0,555 1.60 ok82 1917 0,638 0,562 0,638 0.362 1.00 1.00

TABLE XV SMOOTHED PHASE EQUILIBRIUM DATA FOB THE METHANE-ISOPENTANE BINARY SYSTEM AT 280OF Vapor Phase Composition Liquid Phase Composition Equilibrium Ratio Pressure (mole fraction) (mole fraction) (psia) methane i-pentane methane i-pentane methane i-pentane 500 0,515 0,[87 0,090 0.910 5.70 0.535 75 o.0601 0,599 0.160 O8'84o 3.76 O i75 1000 0,636 o,364 0,231 0.769 2,75 0.473 1250 0,645 0.355 0.318 0,682 2.03 0.520 1500 0,591 0,409 0,460 0,540 1,28 0.757 1534 0.539 0,461 0,559 0.461 1.00 1.00

TABLE XVI SMOOTEIED PHASE EQUILIBRIUM DATA FOR THE METHANE -NEOPENTANE BINARY SYSTEM AT 160~: Vapor Phase Composition Liquid Phase Composition Equilibrium Ratio Pressure (mole fraction) (mole fraction) (psia) methane neopentane methane neopentane methane neopentane 300 O,656 O,344 0,079 0,921 8.30 0.374 500 0,756 0o244 0.150 0,850 5,04 0.287 750 0.798 0,202 0,230 0,770 3 47 0.262 1000 0, 818 0o 182 0,311 0o 689 2.63 0.264 1250 0,814 0,186 o,388 0. 612 2.10 0o304 1500 0,o 789 0.211 0. 475 0.525 1 66 O0 402 1755 o0.644 0356 0.644 0.356 1.00 1,00

TABLE XVII SMOOTHED PHASE EQUILIBRIUM DATA FOR THE METHAIE-NEOPE1NiTANE BINARY SYSTEM AT 2200F Vapor Phase Composition Liquid Phase Composition Equilibrium Ratio Pressure (mole fraction) (mole fraction) (psia) methane neopentane methane neopentane methane neopentane 510 0,405 0,595 0,051 0,949 7,96 0.626 500 0,561 0.439 0,116 O.884 4,84 0,o97 750 o,641 0,559 0,198 0802 3.24 0.448 1000 0o670 0,330 0,281 0.719 2.58 0.459 1250 0,654 0.346 0,574 0.626 1.75 0,555 1460 0,528 0,472 0,528 0,472 1,00 1.00

TABLE XVIII SMOOTED PHASE EQUILIBRIUM DATA FOR THE IETHANE-NEOPENTAINE BINARY SYSTEM AT 280oF Vapor Phase Composition Liquid Phase Composition Equilibriam Ratio Pressure (mole fraction) (mole fraction) (psia) methane neopentane methane neopentane methane neopentane 500 O0 277 O 723 0.o 066 o 934 4.20 0774 750 0o 406 O 594 0 160 o,840 2.54 0.707 1035 0 ~6354 o.646 0,354 o 646 1,00 1.00

2200 -0- THIS WORK — E SAGE ET AL (50) 2000,X 1800 = 220OF / / 1800 / 1600 1 na — 3 O000 i 1400 1 a: I t (I) 0 10 20 30 40 50 60 70 80 90 COMPOSITION (MOLE % METHANE) Figure 4, Pressure-Composition Diagram for Methane-Normal Pentane Binary System at 2200F.

-712400 2200- O THIS WORK E AMICK ET AL ( I ) T = 160 ~F 2000 1800 1600 1400 c 1200 ui cr 0 1000 - ui 400 200 - 600 0 El 400 El El 200 0 10 20 30 40 50 60 70 80 90 COMPOSITION (MOLE %METHANE) Figure 5. Pressure-Composition Diagram for Methane-Isopentane Binary System at 160~F.

-722200 O THIS WORK o AMICK ET AL ( I 2000 T= 220 OF 1800 1600 1400 - 1200 wI 1000 El cr a- 800 600 o 400- El 200 0I I I I I I I 0 10 20 30 40 50 60 70 80 90 -COMPOSITION (MOLE % METHANE) Figure 6. Pressure-Composition Diagram for Methane-Isopentane Binary System at 2200F.

-73 - 2000 1800 0 THIS WORK E] AMICK ET AL (I ) T = 280 F 1600 - 1400 1200 1000 0 0 U) w. 800 _ 600 0 400 -!0 200 O. I I I I I 0 10 20 30 40 50 60 70 80 COMPOSITION (MOLE % METHANE) Figure 7. Pressure-Composition Diagram for Methane-Isopentane Binary System at 280 F.

10.0 8.0 0 160 0F 6.0 - (0 220 F 5.0 - 280 F z 4.0 4 I3.0 2.0 v 1.0 0.8: 0.6 0.5 0.4 - wU Z 0.3 Z a. 0 0.2 0.1 100 200 300 500 700 1000 2000 3000 5000 PRESSURE (PSIA) Figure 8. Equilibrium Ratio-Pressure Diagram for Methane-Isopentane Binary System.,

-75 - 10.0 8.0 - 0 T-1600F, C 0.259 - T=220F, C =0.252 6.0 - A T= 280F, C = 0.257 5.0 4.0 - METHANE 3.0 2.0 - 0 1.0 cr 0.8m 0.6 - 0.5 - ISOPENTANE o n- PENTANE w 0.4 ISOPENTANE 0.1 I I I I I 100 200 400 600 800 1000 2000 5000 PRESSURE (PSIA) Figure 9. Equilibrium Ratio-Pressure Diagram for Methane-Isopentane-Normal Pentane Ternary System.

-762400 1800 1600 _ 40 0 1400 - S 1200. 1000 800 600 400 200 0 I 0 10 20 30 40 50 60 70 80 90 COMPOSITION (MOLE % METHANE) Figure 10. Pressure-Composition Diagram for Methane-Isopentane-Normal Pentane Ternary System at 160~F.

-772200 T=2200F 2000- 0:,. Xi-cM:0.252 1800 0 1600 - 1400 1200 U) a_ 800 - 600 - 0 10 20 30 40 50 60 70 80 90 COMPOSITION (MOLE % METHANE) Figure 11. Pressure-Composition Diagram for Methane-Isopentane-Normal Pentane Ternary System at 220~F.

-782000 T= 2800F 1800 - = 0.257 (o) Xi-c? Xn-c5 C = 0.241 (El ) 1600 1400 1200 - 1000 U) a, cr Z 800 a. COMPOSITION (MOLE %METHANE) Figure 12. Pressure-Composition Diagram for Methane-Isopentane-Normal Pentane Ternary System at 280~F.

-79METHANE 10t 0 90 10 80 -- 20 T= 160 F 30 o 1003 psia o 1493 psia 1995 psia 10 0 50 7 -'0 0 60 4 80 20'90 v$~~~ v v v ~ i00 00 ISOPENTANE MOLEn- PENTANE Figure 13. triangular Composition Diagram for Methane-Isopentane-Norma! Pentane System at 16O~F.

-80METHANE 100 0 90 0 8 20 T 2200F 0 995 psia 70 30 E 1519 psio /\ 1765 psia 100 90 80 7A0d0 4 3 0 _-.. 20 ~~~~/ \~~~~~~~~~~90 10 V2-V v v v v v 100 1oo 90 80 70 60 50 40 30 2d 10 0 ISOPENTANE MOLE % i- C5 n- PENTANE Figure i4. Trianglar Composition Diagram for Methane-Isopentane-Normal Pentane System at 220~F.

METHANE I00OO 90180. 8-~-~ T = Z80 ~F M r \ 2 1031 PSIA ace 0 4 01255 9 0 80 70 60 50 40 30 20 ISOPENTANE M.OLE % iC5 n-PENTANE Figure 15. Triangular Composition Diagram for Methane-Isopentane-Normal Pentane System at 280~F. 2 0 I0 90 0 v. v v v V00 tOO 90 80 70 60 50 40 30 20 10 0 ISOPENTANE MOLE % i-C5 n-PENTANE Figure 15. Triangular Composition Diagram for Methane-Isopentane-Normal Pentane System at 280~F.

-822000 1800 o T = 160 F 1600 1400 1200 LU cr 1000 en) (K) LU aO 800 600 400 200 0 X I I I I I I I I I 0 10 20 30 40 50 60 70 80 90 COMPOSITION (MOLE % METHANE) Figure 16. Pressure-Composition Diagram for Methane-Neopentane Binary System at 160~F.

-83 - 2000 1800 T =220 F 1600 1400 - 1200 - a. 1000 - LI U) 0 800 cr a600 400 - 200 0 0 10 20 30 40 50 60 70 80 COMPOSITION (MOLE % METHANE) Figure 17. Pressure-Composition Diagram for Methane-Neopentane Binary System at 2200~F.

-841600 T= 280 F 1400 1200 1000 mg 800 (1 600 00_ 400 200 0 10 20 30 40 50 60 COMPOSITION (MOLE % METHANE) Figure 18. Pressure-Composition Diagram for Methane-Neopentane Binary System at 280 TF.

-851.0 8.0 -8.LI 0 160~F Z 6.0 - 1: 220~F 5.0 - 4.03.02.0 0 cr.0 D 0.8 r z - W L 0.5 0.. 0 w 0.4 Z 0.3 0.2 0.4 - I00 200 300 500 700 1000 2000 3000 5000 PRESSURE (PSIA) Figure 19. Equilibrium Ratio-Pressure Diagram for Methane-Neopentane Binary System,

-8610.0 8.0 6.0 _ METHANE 5.0 T = 0.247 4.0 3.0 2.0 0 1.0 _ 0.8 o 0.6 - 0.5 - 0.4 0.3 NEOPENTANE 0.2 n- PENTANE 0.1I I I I fOo 200 300 400 600 800 1000 2000 4000 PRESSURE (PSIA) Figure 20. Equilibrium Ratio-Pressure Diagram for Methane-NeopentaneNormal Pentane Ternary System.

2400 2200 T= 160 F 2000 XNEO C5 2000 - 0.247 XNEOC5_+Xn-C5 1800 1600 1400 1200 - a. w cr )n 1000 - w 600 - 400 - 200 - 0 I I I i I I I I 0 10 20 30 40 50 60 70 80 90 COMPOSITION (MOLE %/ METHANE) Figure 21. Pressure-Composition Diagram for Methane-Neopentane-Normal Pentane Ternary System at 160~F.

-88METHANE 100 0 90 10 80 20 70 / 0 0T = 160 F o 1006 psia o 1505 psio 0 60 as 201313pia /70 6 50 40 3 20 o\O 4 2 90 10 IY — 1 00 To 60 50 40 30 n- PENTANE MOLE % NEO-C5 NP iguNTAe 2. Triangular Composition Diagram for 4ethane-NeopentaneN enta n

VIIIo ANALYSIS AND DISCUSSION OF RESULTS The certainty of the experimental results in this research is dependent upon accuracy of measurements, experimental technique, and purity of the materials used~ The measurements made can be divided into three areas: pressure measurement, temperature measurement, and composition determination~ The equilibrium pressures were measured wit'h a Heise pressure gauge. This instrument has a pressure range of 0 to 3,000 psi and its scale is subdivided at 2 pounds per square inch intervals. The model used in this research was accurate to 0,1 percent over its entire range. The gauge was calibrated against a dead-weight tester and found to be accurate over its entire range of full scale. Accordingly, it is believed that the equilibrium pressures were known to within +3 pounds per square inch, Calibration of the Heise pressure gauge is given in Appendix C, Table XXVIII.'he temperature of the constant-temperature bath which contained the equilibrium cell. was determined with a mercury-in-glass thermometer which was calibrated against standard thermometers. The uncertainty in the measurement of temperature is within +,20F. However, temperature variations in the bath were caused by the temperature controller, At temperatures below about 2200F, fluctuations of the bath temperature were below 04~Fo. At a bath temperature of 280~F, the temperature fluctuation was within O,50F. Inasmuch as the equilibrium cell has a relatively large heat capacity, the cell contents would incur smaller temperature variations. It is believed that the overall uncertainty in temperature is + o5oFo Calibration of the mercury-in-glass thermometer is given in Table XXIX of Appendix CO -89

-90The compositions of the dew-point gas and the bubble-point liquid were investigated by withdrawal of a small portion of the vapor and liquid phases under isothermal and essentially isobaric conditions. Normally, at pressures below 0.7 of the critical, about a two to three psi pressure drop occurred after the vapor sample was withdrawn from the cell. At pressures near the critical region, a higher (about 8 psi) drop in pressure occurred upon vapor sample withdrawal. The compositions of the vapor and liquid samples were ascertained by gas chromatography. Duplicate analyses of the same sample were run as standard procedure. The data are tabulated in Appendix B, Tables XXVI and XXVII. The duplicate analyses agreed to within about 0.5 percent of methane for the majority of samples. It is believed that the analytical method is accurate to within 1.0 percent of the main constituent in the mixture. The gas chromatograph was calibrated with binary mixtures of known concentration. Details of the calibration procedure are given in Appendix C. Although the reproducibility of the measurements, in particular the determination of composition, indicates adequate technique in analysis of the vapor phase and liquid phase samples, consideration must also be given to the uncertainty of the experimental technique —that is, certainty of equilibration and the withdrawal of-.representative samples from the cell. By making several runs at the same temperature and essentially the same pressure for the methane-normal pentane binary system, the reproducibility of the system can be determined. The results presented in Table II and Figure 4 indicate a reproducibility of about 1.0 percent.

-91Listed in Table I are the purities of the materials used in this research. Only the methane contains a significant impurity, about 0.7 percent nitrogen, Since the column used in the chromatograph cannot separate nitrogen and methane, the methane compositions determined in this work are slightly biased by the nitrogen impurity. To determine the effect of nitrogen on the experimental results, a material balance was made. Specific volume data reported by Sage et al.(5~) on the methane-normal pentane binary system and K values of nitrogen reported by Roberts and McKetta(47) were used to make a nitrogen material balance. The results of these calculations showed that the maximum nitrogen concentration in the vapor phase was about 1,0 percent, with no significant accumulation of nitrogen in the system due to the sampling technique. This fact is further substantiated when one compares the result of this work on the methane-normal pentane binary system with that of Sage et al, Figure 4 shows the dew-point gas compositions to be slightly greater than those reported by Sage et al, by about 1.0 percent. The liquid phase compositions determined by them, however, are in very good agreement with those of thIis work. Comparison of runs 66 through 69 and runs 66A through 69A further show no significant accumulation of nitrogen due to the sampling technique. Puuns 66 through 69 were the last four runs of a series of eighteen runs made with an initial charge of isopentane-normal pentane mixture, Runs 66A through 69A were the first four runs of a new charge of isopentane-normal pentane mixture, Comparison of the methane concentration in the vapor phase of both sets of runs (see Figure 12) indicates no nitrogen accumulation within the accuracy of the analytical

-92technique. The isopentane to isopentane plus normal pentane concentration in the liquid phase is different, however, for the two sets of runs. This is a result of the relative volatility of isopentane to normal pentane. That is, the concentration on a methane-free basis of isopentane to normal pentane in the liquid phase is somewhat lower for runs 66 through 69 compared to runs 66A through 69A. Figures 5 through 7 present and compare the phase composition data of this work and that reported by Amick, Johnson, and Dodge(1) for the methane-isopentane binary system. The data reported by Amick et al, are not in good agreement with this work. Their data do show considerable scatter, however, as can be seen in Figures 5 through 7. Amick et alo employed a bubble-and-dew-point device to obtain their data. In contrast to their conclusion, the results of this work show the solubility of methane in isopentane not to be very different from the solubility of methane in normal pentane. For the methane-isopentane binary system, critical pressures for the three isotherms were determined graphically by extrapolating to zero. plots of system pressure versus the quantity (y-x)2, The corresponding critical compositions of the binary systems investigated were determined by using the law of rectilinear diameters, where plots of equilibrium pressure versus the quantity (y+x)/2 were extrapolated to the previously determined critical pressures. It is believed that the graphically determined critical pressures have an uncertainty of + 20 pounds per square inch, and the corresponding critical compositions have an uncertainty of less than + 1 mole percent methane, Included in Table XIX are the critical pressures and corresponding critical compositions for temperatures 160 ~F, 220 ~F, and 280 ~F.

-93 - No data have been found in the literature for the methane-neopentane binary system. Therefore, no comparisons are made with data of this work, Critical pressures and critical compositions are indicated in Figures 16 through 18o They were determined in the same manner as described in the discussion of the methane-isopentane binary system. It is of interest to compare the methane-neopentane system with the methane-isopentane binary system, Figures 16 through 18,in contrast with Figures 5 through 7, reveal a significant difference in solubility of methane in these two pentane binary systemso The solubility difference of methane, expressed in terms of the equilibrium vaporization ratio, K in the neopentane solution and the isopentane solution is also borne out by comparing Figures 8 and 10, respectively. One would expect this solubility difference in view of the difference of molecular structure of neopentane and isopentane, Because of its symmetrical structure, it is reasonable to assume that liquid neopentane would contain a larger void fraction than liquid isopentane, On this assumption, one can then visualize an increased solubility of a relatively spherical molecule such as methane TABLE XIX GRAPHICALLY DETERMINED CRITICAL PROPERTIES FOR BINARY SYSTEMS Temperature Critical Pressure Critical Composition System (~F) (psia) (mole fraction methane) methane-isopentane 160 2213 o 688 methane-isopentane 220 1917 o0,638 methane-isopentane 280 1534 0,539 methane-neopentane 160 1755 0o,644 methane-neopentane 220 1460 06528 methane -neopentane 280 1035 0o354 _ _ ~,,~,_~_______, ~_ ~-

-94Illustrated in Figures 9 through 15 is the effect of adding the intermediate constituent isopentane to the methane-normal pentane binary system, Figure 9 reveals that the effect of isopentane does not change the equilibrium vaporization ratio for methane significantly for any of the three isotherms. As can be seen in Figure 9, the K curves have similar characteristic shapes as those for binary systems. Figures 9 through 12 are characterized by an additional intensive property. The chosen property (designated as C ) is the average concentration of isopentane to isopentane plus normal pentane in the liquid phase, An average value was used since this mole fraction did change slightly during an isotherm determination, This is a result of the relative volatility of isopentane to normal pentane, For the methane-isopentane-normal pentane ternary system, the change in the mole fraction parameter for the 160oF and 220~F isotherms was less than 2 percent. For the 2800F isotherm, the change in mole fraction of isopentane to isopentane plus normal pentane was less than 2,5 percent, The triangular compostion diagrams, Figures 13 through 15, show reasonable consistency with the binary data of this work and Sage et al, Figure 20 through 22 present the effect of adding a different intermediate constituent to the methane-normal pentane binary system. For this case, the third constituent is neopentaneo Experimental data were obtained at one isotherm, 1600~Fo The solid lines in Figure 22 are the combining lines connecting the coexisting vapor and liquid phases, Mole fraction of the heavy components on a methane-free basis in the liquid phase is used as the third intensive property in Figures 20 and 21, For the reasons discussed in the preceding paragraph, an average value of

this property is presented in Figures 20 and 21, For this ternary system, the methane- free mole fraction of neopentane to neopentane plus normal pentane incurred about a 3 percent change for the isotherm (1600F) determinat ion. The general behavior portrayed in Figure 20 is similar to that found. for binary systems; however, it does illustrate the influence of neopentane upon the equilibrium behavior of the individual components. Comparison of Figures 9 and 20 indicate a significant influence of the nature of the intermediate component, although a different isomer, upon the phase behavior of the light component, methane. That is, for essentially the same liquid phase parameter, neopentane in contrast with isopentane produces a decrease of the equilibrium ratio of methane for the same temperature and pressure of the system,

IX, ANALYTICAL CORRELATION PROCEDURE The ultimate goal of phase equilibrium thermodynamics is to develop accurate and reliable methods to predict vapor-liquid phase behavior of complex mixtures. However, such methods can only be deemed reliable when subjected to direct comparison with experimental data. The correlation procedure adopted for the calculation of the equilibrium vaporization ratio K for the components investigated in this research is a modified form of the Chao-Seader correlation.(10) Chao and Seader express the equilibrium vaporization' ratio in terms of rigorously defined thermodynamic quantities (Equation (51)) 0 where v. is the liquid phase fugacity coefficient, and is defined as i f~/P, Zi is the liquid phase activity coefficient, and Xi is the vapor phase fugacity coefficient. Equation (51) is obtained by substitution of Equations (18) and (22) into Equation (12). That is, 0x p =, ~(66) Rearranging Equation (66) in terms of the definition of the equilibrium ratio, K. yields Equation (51). In this section a detailed discussion is given on the method employed to calculate the thermodynamic functions in Equation (51), Ao Equation of State A good equation of state is necessary in deriving thermodynamic functions to represent experimental vapor-liquid equilibrium data, For -96

-97this work the equation presented by Benedict, Webb, and Rubin(4) was chosen to evaluate the specific volume of the vapor phase, the compressiblity factor Z, and the vapor phase fugacity coefficient ~ o The reasons for selecting this equation as opposed to other equations of state were twofold. First, the B-W-R equation can be used to predict thermodynamic properties in the critical region. Second, constants used to describe the behavior of the vapor phase region have been determined for all the pure components studied in this research. The relationship between the fugacity coefficient and pressure, temperature, and volume is given by Equation (25). Substitution of the B-W-R equation into Equation (21) yields an expression for the fugacity coefficient 0i: RTm.-= RT., +_ [( o.Bo RT - (o7o) - coi Co ) STY] y + 3 [RT (bar (Ri T )L23] 13 ~+ -.-[(a cacl) +" 1- (a.aa)i/] y5 + (67) +I r~i'i'3[ 4 +? I' -o. -[) 2 (1) [ - _-tPJ_ In the case of this research, the constants used for the computations were obtained from the literature and are given in Table XXo

TABLE XX CONSTANTS FOR BWR EQUATION OF STATE FOR INDIVIDUAL MATERIALS USED IN THIS RESEARCH Substance Methane(4) Neopentane(7) Isopentane(4') Normal Pentane(4) Bo O0.0426000 0.170530 0.160053 0.156751 Ao 1.855500 12.9635 12.7959 12.1794 -6 Co x 106 0.0225700 1.273 1.74632 2.12121 b 0.00338004 0.0668120 O. 0668120 0.0668120 a 0.0494000 3.4905 3.75620 4.07480 c x 10-6 0.00254500 0.546 o0.695000 0.824170 ax 103 0.124359 2.000 1.70000 1.81000 x 102 0.600000 5.000 4.63000 4.75000 Units: P = Normal atmosphere d = gm-moles/liter T = OK = ~C + 273.16 R = 0.08207 (liter)(atm)/(gm-mole)(OK) B. Activity Coefficient The various forms of representing activity coefficients have been previously discussed. Prausnitz, Edmister, and Chao(42) have recommended that the Hildebrand-Scott(20) regular solution theory correlation for activity coefficient be used in non-polar mixtures. Chao and Seader claim good representation of experimental data using this equation.

-99Equation (45a) is given as RT,, 7, 6a [a 6] The solubility parameter designated by the symbol bi is defined as Equation (45b) where E at ordinary temperatures can be identified with the energy where AE at ordinary temperatures can be identified with the energy of vaporation or the energy required to vaporize the liquid to infinite volume, and Vi is the molal volume of constituent i. The symbol - designates the volume average value of the solubility parameter for the solution and is given in mathematical form as: - = L 5z (68) Equations (45a), (45b), and (68) are given by Chao and Seader, At temperatures well below the critical, the energy of vaporization is essentially the enthalpy of vaporization minus the quantity RT, so that Equation (45b) may be rewritten as 8[An XR- 11/a (69) The approximate variation of the solubility parameter with temperature is given by Hildebrand and Scott(2O) as _ 1. 5 (70) where a denotes the coefficient of thermal expansion.

-100Prausnitz, Edmister, and Chao have prepared a plot of log 5 as a function of temperature for several hydrocarbons. Chao and Seader give values of the solubility parameter for one temperature only, since they assume the solubility parameter is independent of temperature. At the outset, values for the solubility parameter used in this correlation were taken from the plot of Praunitz et al. Later constant values for the solubility parameter as given by Chao and Seader were used, Direct comparisons of the predicted equilibrium vaporization ratios, using both sets of values of the solubility parameter, indicated that better agreement between the observed equilibrium vaporization ratios determined from this work and those predicted by the correlation could be obtained when solubility parameter values presented by Chao and Seader were used. The values of solubility parameter used in this work are presented in Table XXI. TABLE XXI SOLUBILITY PARAMETERS 1/2 Component (cal/ml) Source Methane 5. 68 (10) Neopentane 6,25 (20) Isopentane 6.75 (20) Normal Pentane 7.02 (10) Table XXII presents the critical constants used in this correlation, Also included in Table XXII are the values for liquid molal volume and the acentric factor as presented by Chao and Seader.

-101TABLE XXII CONSTANTS FOR PURE COMPONENTS Tc Pc V Component (OF) (PSIA) (ml/gi-mole) Methane -115.8 67,3.1 0.0 52 Neopentane 321.1 464.0 0.195 123.3 Isopentane 370.1 483,0 0.2104 117,4 Normal Pentane 385 9 489.5 0.2387 116,1 C, Fugacity Coefficient of the Pure Liquid Component The fugacity coefficient for a pure liquid is defined as the reference fugacity f0 divided by the total pressure P. The referi ence fugacity of a pure liquid is usually, but not always, taken to be the fugacity of the pure component at a system temperature and under its own vapor pressure. In equation form, we write the reference fugacity as 0 ~ -p D[{] (71) where the term in parenthesis is the liquid-phase fugacity coefficient based on the vapor pressure~ The fugacity coefficient at the vapor pressure can be corrected (to the system pressure by the Poynting effect, Then dividing Equation (71) by the total pressure P gives 0 P~ p~T f ])DO kzdD (72) P p p" Rr~~~~~~~~~~7

-102If integration is carried out assuming an average value for the specific volume for the liquid, the following results 0,.~ P _ ~....] - ~. (P-P ) (73) P RT Generalized fugacity coefficients as functions of reduced temperature, reduced pressure, and critical compressibility factor, Zc' have been presented by Lydersen, Greenkorn, and Hougen.(34) Curl and Pitzer(l5) present generalized fugacity coefficients as a function of reduced temperature, reduced pressure, and a third parameter which they call the acentric factor, A correlation does exist between the critical compressiblity factor and the acentric factor, Riedel(46) demonstrates this relationship. The acentric factor which Curl and Pitzer define as (Equation (31)) -) L 1i. ooo -+9-i JT=o.7 indicates the deviation of the behavior of substances from that of simple fluids, Chao and Seader extend the Curl and Pitzer correlation of the liquid phase fugacity coefficient to conditions where a liquid mixture component does not exist as a pure liquid. The extension is achieved through calculation from experimental vapor-liquid equilibrium data. The analytical expression given by Chao and Seader for the liquid phase fugacity coefficient is

-103 - The term v(o) is the fugacity coefficient of simple fluids which are characterized by a zero acentric factor. The term v(l) is a correction term, and accounts for the departure of the properties of real fluids from those of simple fluidso Chao and Seader have expressed the quantities v () and v(l) as functions of reduced temperature and pressure. These' terms have been fitted with the following functional formso ~-4 UJ 4k - / +. -A4T+, T4 3 (74b) and 2= =-4. 23893 + 8.6 58087T (74c) -1.22060 -3.15224T2, -0-.025[P -06] The coefficients in Equation (74b) as given by Chao and Seader are presented in Table XXIIIo TABLE XXIII CONSTANTS FOR LIQUID PHASE FUGACITY COEFFICIENT EXPRESSION Simple Fluid Methane Ao 5075748 2,43840 A1 -3 01761 -2o,24550 A2 -4, 98500 -0o34084 A3 2, 02299 0000212 A4 0 -0.00223 A5 0 08427 0,10486 A6 0, 26667 -0o o3691 A -0, 311 8 o A8 -0~02655 0 A9 0, 02883 0

-104The acentric factors' listed in Table XXII are not derived from the original definition. The values are those presented by Chao and Seader and were determined as a parameter for the best fitting of the vapor pressure data for pure components according to the Chao-Seader correlation given by Z: 6 4 8L[ t U> ] (75) A computer program was written to calculate equilibrium vaporization ratios from Equation (51). At the outset Equations (45), (67), and (74) were used in conjunction with Equation (57) to calculate the equilibrium vaporization ratios of the compounds studied in this work. After several trial runs, it became apparent that the calculated equilibrium ratios for methane were always greater than observed values, and that this discrepancy increased with increasing pressure. In view of the fact that the formulation of activity coefficients (Equation (45a)) is independent of pressure and that the Benedict, Webb, and Rubin equa(4) tion is believed to be reliable in representing the P-V-T behavior of gaseous mixtures, the expression for liquid phase fugacity coefficient was modifiedo Since the acentric factor equals zero for methane, Equation (74a) reduces to 0 (=- o)(76) The expression for log v(o) (Equation (74b)) was then divided by the quantity (l+PxlO-4) where P is the total pressure of the system. It should be remembered that this quantity has no theoretical implications and that it is only a first order approximation to better fit the experimental data.

Table XXV in Appendix A presents comparisons of the calculated equilibrium vaporization ratios with the observed equilibrium vaporization ratios. The percent deviation is defined by Percent Deviation KOBSx 100 (77) KoBs Also included in this table are the numerical values for the vapor and liquid phase fugacity coefficients, liquid activity coefficients, vapor specific volume, and the compressiblity of the vapor. At the end of each system investigated in this research is the numerical value for the average absolute percent deviation. As can be seen, calculated K values and observed K values are in reasonably good agreement, except near the critical region. The average absolute deviation for the methane-normal pentane binary system is about 4 percent. For the methane-isopentane binary system, the average absolute deviation is within 8 percent. The average absolute percent deviation for the methane-neopentane-normal pentane ternary and methane-isopentanenormal pentane ternary systems are 6.3 percent and 6.0 percent, respectively. Inspection of Table XXV for the methane-neopentane binary system reveals the predicted methane K values to be in greater error than the predicted neopentane K values, especially at higher temperatures. This observation concurs with the temperature restriction for methane imposed on the Chao-Seader correlation. In other words, the Hildebrand equation cannot predict accurately methane behavior at temperatures above 0.93 of the pseudocritical temperature of the equilibrium liquid mixture. It appears unlikely that the complex behavior of liquid

mixtures composed of constituents with such different physical properties as methane and pentanes can be represented in the critical region by such relatively simplified expressions as Equations (45) and (74).

X. SUMMARY AND CONCLUSIONS The methane-normal pentane binary system was investigated at a temperature of 2200F. Comparison of this work with previous investigations is good. Phase equilibrium data were obtained for the methane-isopentane binary system at temperatures of 1600F, 2200F, and 280~F and pressures from about 500 pounds per square inch up to the critical region, The data are presented in both graphical and tabular form. Vapor-liquid equilibrium data have been obtained throughout the coexisting-phase region for the methane-neopentane binary system at pressures from about 300 pounds per square inch to the critical region for temperatures of 160F, 220~F, and 2800~F Pressure versus composition curves and equilibrium vaporization ratio versus pressure diagrams are included, Experimental equilibrium vaporization ratios of methane are significantly lower in a methane-neopentane binary system than in a methane-isopentane binary system. The critical pressures are lower for the methane-neopentane binary system compared to the methane-isopentane binary system for temperatures of 160~F, 2200F and 2800Fo Equilibrium vaporization ratios have been experimentally determined for the methane-isopentane -normal pentane ternary system at temperatures of 1600F, 220~F, and 280OF and pressures from about 500 pounds per square inch up to the critical region, Data are presented graphically eand are also tabulated, The two-phase equilibrium behavior for the methane-neopentanenormal pentane ternary system has been experimentally determined for a -107

temperature of 1600F and pressures from about 500 pounds per square inch up to the critical region. The experimental results are tabulated and presented graphically. A computer program has been written to calculate equilibrium vaporization ratios. The correlation employs the Benedict, Webb, and Rubin(4) equation of state to predict vapor phase fugacity coefficients. Hildebrand's regular solution theory is applied to the liquid phase. And the expression given by Chao and Seader(lo) is used to calculate the liquid phase fugacity coefficient. Comparison of the K-value correlation with all the experimental points determined in this work indicate an average absolute percent deviation within 10 percent,

REFERENCES 1. Amick, E.H., Jr., Johnson, W.B., and Dodge, B.F., "P-V-T-X Relationships for the System: Methane-Isopentane," Chemical Engineering Progress Symposium Series, 48, 3, 65 (1952). 2. Beattie, J.A., and Bridgeman, O.C., "A New Equation of State for Fluids," Proc. Am. Acad. Arts & Sci., 63, 229 (1928). 3. Beattie, J.A., Douslin, D.R., and Levine, S.W., "Vapor Pressure and Critical Constants of Neopentane," J. Chem. Phys., 19, 948 (1951). 4. Benedict, M., Webb, G.B., and Rubin, L.C., "An Empirical Equation for Thermodynamic Properties of Light Hydrocarbons and Their Mixtures," Chem. Eng. Progress, 47, 8, 419 and 517 and 609 (1951); J. Chem. Phys., 8, 334 (1940); J. Chem. Phys. 10, 747 (1942). 5. Boomer, E.H., Johnson, C.A., and Piercey, H.G.A., "Equilibria in Two-Phase Gas-Liquid Hydrocarbon Systems II. Methane and Pentane," Can. J. Research, B16, 319 (1938). 6. Brainard, A.J., "A Study of the Vapor-Liquid Equilibrium for the Quaternary System Hydrogen-Benzene-Cyclohexane-Normal Hexane," Doctoral Thesis, The University of Michigan, Ann Arbor, 1964. 7. Canjar, L.N., Smith, R.F., Volianitis, E., Galluzzo, J.F., and Cabarcos, M., "Correlation of Constants in the Benedict-Webb-Rubin Equation of State," Ind. & Eng. Chem., 47, 6, 1028 (1955). 8. Carlson, H.C., and Colburn, A.P., "Vapor-Liquid Equilibria of NonIdeal Solutions," Ind. & Eng. Chem., 34, 581 (1942). 9. Case, L.0., Elements of the Phase Rule, Edwards Letter Shop, Ann Arbor, Michigan, 1939. 10. Chao, K.C., and Seader,J.D., "A General Correlation of VaporLiquid Equilibria in Hydrocarbon Mixtures," A.I.Ch.E. Jour., 7, 4, 598 (1961). 11. Chueh, P.L., Muirbrook, N.K., and Prausnitz, J.M., "Part II. Thermodynamic Analysis," A.I.Ch.E. Jour., 11, 6, 1097 (1965). 12. Chueh, P.L., and Prausnitz, J.M., "Vapor-Liquid Equilibria at High Pressures; Partial Molal Volumes in Multicomponent Liquid Mixtures,l" to be submitted to A.I.Ch.E. Jour. 13. Curl, R.F., Jr., and Pitzer, K.S., "Volumetric and Thermodynamic Properties of Fluids —Enthalpy, Free Energy and Entropy," Ind. & Eng. Chem., 50, 265 (1958). -109

-11014. Flory, P.J., "Thermodynamics of High Polymer Solutions," J. Chem. Phys., 10, 51 (1942). 15. Gibbs, J.W., The Collected Works of J.W. Gibbs, Volumes I and II Longmans, Green and Company, New York, 1931. 16. Grayson, H.G.,, and Streed, CoW., "Vapor-Liquid Equilibria for High Temperature, High Pressure Hydrogen-Hydrocarbon Systems," presented at the Sixth World Petroleum Congress in Frankfurt/Main, June, 1963. 17. Hadden, S.T., "Convergence Pressure in Hydrocarbon Vapor-Liquid Equilibria," Chem. Eng. Progress Symposium Series, 49, 7, 53 (1953). 18. Heichelheim, H.R., Kobe, K.A., Silberberg, I.Ho,, and McKetta, J.J., "Compressibility Factors of 2,2-Dimethylpropane (Neopentane)," J. Chem. Eng. Data, 7, 4, 507, (1962). 19. Hildebrand, J.H,, and Scott, R.L., Regular Solutions, Prentice-Hall, Inc., Englewood Cliffs, New Jersey, 1962. 20. Hildebrand, J. Ho,, and Scott, R.L., The Solubility of Nonelectrolytes, Dover Publications, Inc., New York, 1964. 21. Hildebrand, JoH., and Wood, S.E., "The Derivation of Equations for Regular Solutions," J. Chem. Phys., 1, 817 (1933). 22. Hirschfelder, J00., Curtiss, CF., and Bird, R.B., Molecular Theory of Gases and Liquids, John Wiley and Sons, Inc., New York, 1954. 23. Huggins, M.L., "Thermodynamic Properties of Solutions of Long-Chain -Compounds," Annals of New York Academy of Science, 43, 1 (1942). 240 Katz, D.L., Cornell, D., Kobayashi, R., Poettman, F.H., Vary, JoAo,, Elenbaas, JoRo, and Weinaug, C.F., Handbook of Natural Gas Engineering, McGraw-Hill Book Company, Inc., New York, 1959. 25. Katz, D.L., and Hachmuth, K.H., "Vaporization Equilibrium Constants in a Crude Oil-Natural Gas System," Ind. & Eng. Chem, 29, 1072 (1937). 26. Katz, DoL., and Kurata, F0, "Retrograde Condensation," Ind & Eng. Chem., 32, 817 (1940). 270 Kellogg Charts, Liquid-Vapor Equilibrium in Mixtures of Light Hydrocarbons, M. Wo Kellogg Company, Inc. New York, 1950. 28. Kihara, T., "Virial Coefficients and Models of Molecules in Gases," Revo Modern Phys., 25, 4, 831 (1953). 29. Lennard-Jones, JoE, "On the Determination of Molecular Fields-II from the Equation of State of a Gas," Proc. Royal Soc. (London), A106, 463 (1924).

-11130. Lenoir, J.M., and White, G.A., "Predicting Convergence Pressures," Petroleum Refiner, 37, 3, 173 (1958). 31. Lewis, W.K., and Luke, C.D., "Properties of Hydrocarbon Mixtures at High Pressures," Trans. ASME, 54, 17, (1932). 32. Lewis, G.N., and Randall, M., Thermodynamics, McGraw-Hill Book Company, Inc., New York, 1923. 33. Lyckman, E.W., Eckert, C.A., and Prausnitz, J.M., "Generalized Liquid Volumes and Solubility Parameters for Regular Solution Application," Chem. Eng. Sci., 20, 703, (1965). 34. Lydersen, A.L., Greenkorn, R.A., and Hougen,O.A., Engineering Experiment Station Report Number 4, University of Wisconsin, Madison, 1955. 35. Martin, J.J., and Hou, C.Y., "Development of an Equation of State for Gases," A.I.Ch.E. Jour., 1, 142 (1955). 36. NGAA K-Value Charts, Engineering Data Book, Natural Gasoline Supply Men's Association, Tulsa, Oklahoma, 1957. 37. O'Connell, J.P., and Prausnitz, J.M., "Thermodynamics of Gas Solubility in Mixed Solvents," Ind. & Eng. Chem. Fundamentals, 3, 4, 347 (1964). 37a. Organick, E.I., "Equilibrium Ratio Charts for Hydrocarbon Systems, a new Publication by NGAA," Proceedings Thirty-Fourth Annual Convention, Natural Gasoline Association of America, April, 1955. 38. Orye, R.V., and Prausnitz, J.M., "Multicomponent Equilibrium with the Wilson Equation," Ind. & Eng. Chem., 57, 18 (1965). 39. Pitzer, K.S., Lippman, D.Z., Curl, R.F., Jr., Huggins, C.M., and Petersen, D.E., "The Volumetric and Thermodynamic Properties of' Fluids. II," J. Am. Chem. Soc., 77, 3433 (1955). 40. Prausnitz, J.M., "Thermodynamic Representation of High-Pressure Vapor-Liquid Equilibria," Chem. Eng. Sci., 18, 613 (1963). 41. Prausnitz, J.M., Eckert, C.A., Orye, R.V., and O'Connell, J.P., Computer Calculations for Multicomponent Vapor-Liquid Equilibria, Prentice-Hall, Inc., Englewood Cliffs, New Jersey, 1967. 42. Prausnitz, J.M., Edmister, W.C., and Chao, K.C., "Hydrocarbon VaporLiquid Equilibria and Solubility Parameter," A. I.Ch.E. Jour., 6, 2, 214, (1960). 43. Redlich, 0., Ackerman, F.J., Gunn, R.D., Jacobson, M., and Lau, S., "Thermodynamics of Solutions," Ind. & Eng. Chem. Fundamentals, 4, 4, 369 (1965) 44. Redlich, 0., and Dunlop, A.K., "Thermodynamics of Solution: VIII. An Improved Equation of State," Chem. Eng. Progress Symposium Series, 59, 44, 95 (1963)

-11245. Redlich, 0., and Kwong, J.N.S., "On the Thermodynamics of Solutions. V, " Chem. Reviews, 44, 1, 233 (1949). 46. Riedel, L., "Kritischer Koeffizient, Dichte des Gesattigten Damfes und Verdampfungswarme, " Chem. Ing. Tech., 26, 679 (1954). 47~ Roberts, L.R., and McKetta, J.J., "Vapor-Liquid Equilibria in the n-Butane-Methane-Nitrogen System," J. Chem. Eng. Data, 8, 161 (1963). 48. Rzasa, M.J., Glass, EoEo,, and Opfell, J.B., "Prediction of Critical Properties and Equilibrium Vaporization Constants for Complex Hydrocarbon Systems," Chem. Engo Progress Symposium Series, 48, 2, 28 (1952). 49o Sage, B.H,, and Reamer, H.H,, "Some Methods of Experimental Study of Vapor-Liquid Equilibria," Chemical Engineering Progress Symposium Series, 48, 2, 3 (1952). 50. Sage, B.H., Reamer, HH., Olds, R.H., and Lacey, WoNo, "Phase Equilibria in Hydrocarbon Systems," Ind. & Eng. Chem., 34, 9, 1108 (1942)o 51o Scatchard, G., "Equilibria in Non-Electrolyte Solutions in Relation to the Vapor Pressures and Densities of the Components," Chem. Rev., 8, 321 (1931). 52, Scatchard, G., "Change of Volume on Mixing and the Equations for Non-Electrolyte Mixtures, " Trans. Faraday Soc., 33, 160 (1937). 530 Scatchard, G., and Hamer, WJ., "The Application of Equations for the Chemical Potentials to Partly Miscible Solutions," J. Am. Chemo Soc., 57, 1805 (1935). 540 Souders, M., Jr., Selheimer, W.W., and Brown, G.G., "Equilibria Between Liquid and Vapor Solutions of Paraffin Hydrocarbons," Ind, & Engo Chem., 24, 517 (1932). 54a. Stalkup, F.I., and Kobayashi, R., A.I.Ch.E. Jour., 9, 121, (1963). 55. Wilson, G.M., "A New Expression for the Excess Free Energy of Mixing," J. Am. Chem, Soc., 86, 127 (1964)o 56, Winn, FoW., "Simplified Nomographic Presentation —Hydrocarbon VaporLiquid Equilibria," Chemo Eng. Progress Symposium Series, 48, 2, 121 (1952). 57. WohL, K., "Thermodynamic Evaluation of Binary and Ternary Liquid Systems," Trans. A. I. Cho E., 42, 215 (1946).

APPEND IX A CORRELATION OF VAPOR-LIQUID EQUILIBRIUM DATA Table XXIV presents the computer program written in the MAD (Michigan Algorithm Decoder) language to calculate equilibrium vaporization ratios of the compounds studied in this research, The program is divided into four sections, In the first section, the vapor phase fugacity coefficient is calculated using the BWR equation of state. In the second section use is made of Hildebrand's regular solution theory to calculate the liquid activity coefficient. In the third section of the program, the expression given by Chao and Seader is used to calculate the liquid phase fugacity coefficient. Finally, in the fourth section, comparisons are made between the observed and calculated equilibrium vaporization ratios. Table XXV presents the results of the analytical model used to predict the phase behavior of the components studied in this research. -ll3 -

TABLE XXIV MAD COMPUTER PROGRAM FOR EQUILIBRIUM VAPORIZATION RATIO CALCULATION

-115$ COMPILE MAD, EXECUTF,P. INT OBJECT If)lJ'U1P MAD (06 JAN 1967 VERSIC;N) PROGrRAM LISTING......... BOOLEAN B1O0'L,SWITCH DIHMEINSIuN PC(q4), TC(4), OMEG-(4), VBAR(4), LNt.l(4), 4 NU(4), LNUO(4), 1 LNUI(4), SP(4), Y(4), YCALC(4), X(4), XCALC(4), AO(4), 2 RO(4), Cn(4), B(4) A(4), C(4) AL.F(4), GAM(4), PR(4), 3 TR(4) RiTLNPH(4)) LNPH (4), PHI(4 ), KEXP(4),KCALC (4),TF(4), 4 PR(4),LNSP(4) AC(4),PCC (4),TCC 4) KDEV(4) INTEGFR, J,N,( CUOJNT, iSU,\l NS UT,RUNUM,I ND CALCULATION O()- C':NISTANTS FOR THE BWR EQUIATION START REFAD DATA NrPC,TC,OMEG,V3ARk,AO,BD,CO,A,R,GAM,ALF tNSUPH WHENVER. N.E. 2 WHENEVER N;SJU. F. 1 PIr.INT C(iNM4FNT$llFTHANE NORMAL PENTANF RINARY SYSTFM$ K w.H[t.NFVFfR NSUB. F. 2 PRF [NT (,J MMIFNIT.1HMFETHAN!F ISOUPFNTAN4E BINARY SYSTEMS.TH I:RW I S F PR I NT Ci1M-MENT$l 1 EfTHANE NFOPENTANF tI1 NARY SYSTEMS$ ENi) CIf ([ I T ITIONAL OTHERW I SF WHENEVER NSUBT.E. 1 PRINT CUM!-M T$lMETHA NE ISOPENTANE NORMAL PENTANE TERNARY SYST 1 E.N$ (111- RWT SE PRINT COMMENT $IMFTfHANE NEOPENTANE NORMAL PENTANE TFRNARY SYS 1 TEM$ END) )F COINDITIO)NAL FN[) OF CONDITIONAL AIDEV = 0.0 IND) = 0 C(UN1FJT = 0 t3 O()L = 1 P BEGIN ifEA) D)ATA PRESS,TF,Y,X,SP, BOOL RFAI) F'ORMAT QOOOO,RlJNUM VE[CTOR VALUES 00QO = $C6*$ IND = IND + I TtHROUGH LPO, FOR I = 1,1, I.G. N LPO KFXP(I) = Y(I)/X(I) AOMIX = c).C THRKCOUGC; LP1, FUR. I = l 1, I.G.N SUM = Y( I)*SQPT. (AO( I )) LP1 A.M I X = ACM I X + SUM A(3MIX = A1MIX.P. 2 BOMIX = 0.0 THROUlGH LP2, FOR I = 1,1, I.G.N SUM = Y(I) *BO,(I) LP2 HBMIX = RBO I X + SJUM COMIX = 0.0 THROUGH LP:R3, FOR I = 1,1, I.G.N SUM = Y(I)* SQRT.(CO(I)) LP3 COMIX = SUM + C(oMIX COMIX = COMIX.P.2 BMIX = 0.0

-116K = 1./3. THROUGH LP4, FOR 1=1,1, I.G.N SUM = Y( I)*B(I)P.K LP4 BMIX = BMIX + SUM BMIX = BMIX.P.3 AMIX = 0.0 THROUGH LP9, FOR I =1,1, I.G.N SUM = Y(I) A(I).P.K LP5 AMIX = AMIX + SUM AMIX = AMIX.P.3 CMIX = 0.0 THROUGH LP6, FOR I=1, 1,t I.G.N SUM = Y(I) *C(I).P.K LP6 CMIX = CMIX + SUM CMIX = CMIX.P.3 ALFMIX = 0.0 THRIUGH LP7, FOR I=191, I.G.N SUM = Y(I) * ALF(I ).P.K LP7 ALFMIX = ALFMIX + SUM ALFMIX = ALFMIX.P.3 GAMMIX = 0.0 THROUGH LP8, FOR I= 1,1, I.G. N SUM = Y(I) * SQRT.(GAM(I)) LP8 GAMMIX = GAMMIX + SUM GAMMIX = GAMMIX.P.2 CALCULATION OF SPECIFIC V(OLUME BY BWR EQUATION IlF STATE PRESS = PRESS/ 14.696 PRESA = PRESS*14.696 R = 0.0820544 ITER T = (TF + 459.67)/ 1.8 RT = R*T PGIVFN = PRESS V1=R T/PGIVEN DELV = 0.I*V1 K1=BOMIX*R T-AOMIX-COMIX/(T*T) K2=H MIX*R T-AMIX'+AMIX*ALFMIX/ (V*VI*V ) K3=(CMIX/( TT ))*( O+GAMM I X/(VI*V1) )*(EXP. {GAMMI X/1 (V*V1) ) ) PI=R T/VI+K 1/ ( Vl*V )+ (K2+K3)/(Vl*Vl*V. ) QSIN = P1 - PRESS S1 VI = VI - DELV K=BOM I X*R T-AOMIX-COMI X/(T*T) K2=BMI X*R T-AMI X+AMIX*ALFMIX/( VL*VI*Vl) K3=(CMIX/(T*T))*(1.O+GAMMIX/(Vl*Vl))*(EXP.(GAMMIX/1 (Vl*Vl))) Pl=R T/Vl+K1/(VI*Vl )+ (K2+K3) / (V*V*V1) PSIN = P1 - PRESS WHENEVER PSIN*QSIN.G. 0.0 TRANSFER TO S1 OR WHENEVER PSIN*QSIN.L. 0.0 V1 = VI + DELV DELV = DELV/2. WHENEVER.ABS. PSIN.L..10, TRANSFER TO ROOT TRANSFER TO S1 OTHERWI SE TRANSFER TO ROOT END OF CONDITIONAL ROOT CONTINUE

-117CALCULATION OFn THE VAPOR FUGACITY COEFFICIENT D3 = 1i./V Z = PRESS*V1/RT THROUGH LP9, FOR I = 1,1, I.G. N RTLNPH(I)= R.T*ELOG.( 1./Z) + ( (BOMIX + i30( I ))*RT - 2.*SQRT. 1 (AOMIX*AO(I)) - 2.*SQRT.(COM[IX*CO(I))/T/T)*D3 + 3./2.*(RT*( 2 BMIX.P.2B ( I ) ).P.K - (AMI X.P.2*A( I) ).P.K) ) *03.P.. 2 + 3./5. 3 (AMIX*(ALFMIX.P.2 *ALF( I)).P. K + ALFMI X*(AMIX.P.2 *A(I)).P.K) 4 *D3.P.5 + 3.*D3.P.2*(CMIX.P.2*C(l)).P.K/ T/T -(1. - EXP.( 5 -GAMMIX0D3. P. 2) )/IGAMM[XI*D3.P.2)-EXP. (-GAMMIX*')3. P.2)/2.) 7 -2.01)3.P.2'CMIX/T/T*SQRT. (GAM(I)/GAMMIX)*, ((1.-EXP. (-GAMMIX 8 *D3.P.2))/(GAMMIX*D3.P.2) -EXP. ( -GAMMIX D3.P.2) -GAMMI X* 8 D3.P.92 FXP.(-GAMMIX*D3.P.2)/2.) LNPHI ) = RT LNPH( I)/RT LP9 PHI(I) = EXP.(LNPH(I)) CALCULAT InN OF LIQUID PHASE FUGACITY COEFFICIENT THROU(GH LP10, FOR I=1,1, I.G.N TC(I)=(TCC( )+459.69)/1.8 PC (I) =PCC ( I / 14.696 TR(I) = T/TC(I) PR(I) = PRESS/PC(I) LNIl1( I)= -4.23893 + 8.6580(8*TR(I) -1.2206/TR( - 3.15224* 1 TR ( I ). P.3 -0.()25( PR ( I ) - 0.6) WHENEVER I. F.1 LNUO(() =(2.4384 -2.2455/TR(I) -0.34084TR(I ) +0.00212* 1 TR(I).P.2 -O.00223*TR(I).P.3 +(0.10486-0.03691*TR(I) 2 )*PR(I) - ELOG.(PR(I))/2.303)/(L.+PRFSA*1.E-4) OTHERWISE LNUJ( I ) = 5.75748 -3.01761/TR ( I ) -4.985*TR ( I ) +2.02299*TR ( ) 1.P.2 +(0.08427 + 0.26667*TR(I) -0.31138*TR(I).P.2)*PR(I) + 2 (-0.02655 + 0.02883* TR I) )PR( I ). P.2 - ELOG.(PR( I ) )/2.303 ENO CF CONDITIONAL LNU( I )=LNU( I ) +OMEG (I) * LNU1 ( I) LNU(I) = 2.303*L NULI ) LP10 NU(I) - EXP.(LNJ(I )) CALCULATION OF ACTIVITY COEFFICIENT RAC= 1. 987 RTAC= RAC * T NUM = 0.0 DEN = 0.0 THROUGH LPll1, FOR I=1,1, I.G.N SNUM, = X(T) * VAR( I ) * SP I ) NUM=NUM + SNUM SDEN = X(I) * VBAR(I) LP11 DEN = DEN + SDEN SP83 = NUM/DEN THROUGH LP12, FOR 1= 1,1, I.G.N LNSP I ) = VBAR(I) * ( SP(I)-SPB).P.2/RTAC LP12 ACcI) = EXP.(LNSP(I)) THROUGH LP13, FOR I = 1,1, 1.G. N KCALC (I) = AC(I) * NU (I) / PHI(I) KDEV(I)= IKEXP( I )-KCALC(I) )100./KEXP(I) COUNT = COUNT + 1 LP13 AOEV = ADEV +.ABS.SKDEV(I))

-118WHENEVER IND.E. 4 PRI NT FORMAT QQ1,RUNUM VECTOR VALUESQQ1=$1H1,H*RUN NUMBER *,C6*$ IND = 1 OTHERWISE PRINT FORMAT QQ1A, RUNUM VECTOR VALUES QQlA=$1H-,H*RUN NUMBER *,C6*$ END OF CONDITIONAL PRINT FORMAT QQ2,PRESA,TF VECTOR VALUES(Q2=$1HO,H*PRESSURE = *,F5.0,H* PSIA*,S5,H*TEMPERAATURE = 1,F4.0,H* DEG F**$ PRINT FORMATQQ2A,V1,Z VECTOR VALUESQQ?2A=$1H,H*VAP. SPEC. VOL. = *,F6.4,H* LIT/GMOLF*,S5, 1 H*COMPRESSIBILITY = *,F4.3*$ WHENEVER N.E.2 WHENEVER NSUB.F.1 PRINT FORMAT QQ3 VECTOR VALUES QQ3=$1HO,S25,H*METHANE NOR-PENTANE**$ OR WHENEVER NSUB.E. 2 PRINT FORMAT OQ4 VECTOR VALUESQQ4=$1HO,S25,H*METHANE ISOPENTANE**$ OTHERWISE PRINT FURMAT QQ5 VECTOR VALt ES(QQ0 =$ 1HO, S 2, H*METHANE NEOPFNTANE**$ END OF CONDITIONAL PRINT FORMAT (Q8,Y(I),Y(2) VECTOR V.ALtJES Q(,)8=$1H,H*VAPO1R PHASE COMP*,S)O,f5.4, S12,FS.4*$ PRINT FORMAT QQ9,PHI (1),PHI(2) VECTOR VALUESOQ9 =$1H,H*VAP[]R PHASE FUG C[EF*,S5,F6.4,Sll,F6.4*$ PR I NT FORMATQQ 1O, X ( 1 ), X(2) VECTORK VALUESO010=$IHH,t*LIQID PHASE CCMP*,S9,F5.4, S12,F5.4*$ PRINT FORMATQ.Q11,NU(1),NU[(2) VFCTlR VALIJES(QOl=$1H,H*LIOUID PHASE FU)G COEF*,S3,F6..3,Sll,F6.3*$ PRINT. FC(RMATQQ22, AC(1), AC(2) VECTOR VALUESQ(Q2=$IH,H*L I QI ) ACT COEF*,S1O, F.3,S 2,FS.3*$ PRINT Fl:KMATQ 1 3,KEXP( 1 ),KEXP(2) VFCTOR VALUFSQQ13=$1H,H*K OBS*,S19,F6.3,Sll,F6..3*$ PRINT FURMATQQ 14,KCALC ( 1 ), KCALC (2) VECTCOR' VALUESQQ14=$i.H,H*K CALC*,S18,F6.3,S11,F6.3*$ PRINT FLJRMATQQ15,KDLV( 1),KDEV(2) VFCTOR VALUES0g1S=$IH,H*PERCENr [)FV*,S 1,F7.2,S 10,F7.2*$ OTHERWI SE WHENEVER NSUBT. E.1 PRINT FORMAT Q06 VFCTOR VALUES QQ6,=$1HO,S25,H*METHANE ISOPENTANE NOR-PENTA 1 NE**$ OTHERWISE PRINT FORMAT QQ7 VECTOR VALUESQQ7=$1HO, S25,H*METHANE NEOPENTANE NUR-PENTAN 1 E**$ END OF CONDITIONAL PRINT FORMATQQ16,Y( I ), Y( 2 ) Y(3) VECTOR VALUFSQQ16=$IH,H*VAPOR PHASE COMP*,S10,F5.4tS12,F5.4,S13,F5.4*$ PRINT FCO.RMAT QQ17,PHI(1),PHi(2),PHI(3) VECTOR VALUESQQ17=$lH,H*VAPOR PHASE FUG COEF*, S5F6.4,S11F6.4, S 12, 1 F6.4*$ PRINT FORMATQO18,X( 1), X( 2), X( 3) VFCTOR VAIUESQQ18=$1H,H*LIQUID PHASE COMP*,S9,FS.4,S12,F5.4 S13, F5.4*$ PRINT F]RMATQQ19,NU(1),NU(2)3,NU(3) VECTOR VALUESQQ19=$1H,H*L I QUI D PHASE FUG COEF*,S3,F6. 3, S 11,F6.3, S12,

J F6.3*$ PR I NIT F i-RMAT(2OAC (I),AC (2),AC (3) VECTORe VALUFSQQ2O=$IH,H*-L IQUID' ACT COEF*,S1OF5.3,S 12,F5.3,S13,E5. 3*$ PR INT fk r'IAIQQ)21LKEXP(I ),9KEXP(2),KEXP(3) VFCTOk VALUlFS)Q21=$IH, GBS4 S 1 9,F6.3,S11,F6.3,Sl3, F5.3*$ PRINT F0RMATQQ22,KCALC(flKCALC(2),KCALC(3) VFCTCF VALUL ES - 22=-$l$H,H*K CALC'3,S18,F6. 3,S11,F6.3,S1l2,F6.3*$ PRINT F JAkFMATQQ)23,KDEV(I),KDEV(2),KDEV(3) VFCTUR VALU%0 Q23=$IH,H'PERC[NT DEV*, SII El,F7.2, SlOF7.2, S I F7.2*$ END O(F COFI CLT ITIUNAL WHENFVF'" ROOL.E. 1B TRANSFER TU B[E3GIN OT H ERW WISF lADF V = ADFV/COIJNT PkINT FP4RkMAT,050,gTADFV V E CTU VA"LU\1 FS)Q,90= IH O,H*.AVE. ABS. PERCENT DEV. FUR SYSTEM =*IF6.3* TRANSFER TO START FND OF CONIDITInN AL EN GF P JGRAM.. THE FOLLOWING NAMES HAVE f(CCUIRRED ONLY ONCIE IN THIS PR.OGRAM. COMPILATION WILL CONTINUE. ITER C*7 3

-120TABLE XXV COMPARISON OF OBSERVED AND CALCULATED VAPOR-LIQUID EQUILIBRIUM DATA

METHANE NORMAL PENTANE BINARY SYSTEM RUN NUMBER 24 PRESSURE = 1023 PSIA TEMPERATURE = 220 DEG F RUh NUMBER 21 VAP. SPEC. VOL. =.3783 LIT/GMOLE COMPRESSIBILITY.850 PRESSURE = 1502 PSIA TEMPERATURE = 220 DEG F METHANE NOR-PENTANE VAP. SPEC. VOL. =.2454 LIT/GMOLE COMPRESSIBILITY =.809 VAPOR PHASE COMP.8060.1941 VAPOR PHASE FUG COFF 1.0061.4268 METHANE NOR-PENTANE LIQUID PHASE COMP.2530.7470 VAPOR PHASE COMP.8080.1920 LIQUID PHASE FUG COEF 2.762.115 VAPOR PHASE FUG COEF 1.0092.2985 LIQUID ACT COEF 1.098 1.005 LIQUID PHASE COMP.3800.6200 K OBS 3.16.260 LIQUID PHASE FUG CrEF 1.985.092 K CALC 3.016.270 LIQUID ACT COEF 1.080 1.013 PERCENT DEV 5.34 -3.90 K OBS 2.126.310 K CALC 2.123.313 PERCENT DEV.15 -1.12 RUN NUMBER 25 PRESSURE = 1001 PSIA TEMPFRATURE = 220 DEG F RUN NUMBER 27 VAP. SPEC. VOL. =.4094 LIT/GMOLE COMPRESSIBILITY =.900 PRESSURE = 1265 PSIA TFMPERATURE = 2?0 DEG F METHANE NOR-PENTANE VAP. SPEC. VOL. =.29',2 LIT/GMLF C()MPRFSSIBILITY =.831 VAPORt PHASE CUMP.8050.1953 VAPOR PHASE FUG C:iFF.9582.4291 -'AF THA' iN E NORK-PENTANE LIQUID PHASE CJMP.2470.7530 VAPOR PHASE COMP.8110.f185 LIQUJID PHASE FUG C{]EF 2.817.116h VAPOR PHASE FIJU COEF 1.0042.3587 LIQUID ACT C(]EF 1.(99 1.005 LIQUID PHASE CnrP.3240.6760 K OBS 3.259.259 LIQUIC PHASE FUG Cf;EF 2.292.101 K CALC 3.232.272 LIQUID ACT COUF 1.088 1.009 PERCENT DEV.84 -4.94 K OBS 2.503.279 K CALC 2.483.284 PERCENT DEV.81 -1.95 RUN NUMBER 26 PRESS.JE = 19g99 PIA TEMPKl<ATURE = 220 DEG F RUN NUMBER 23 VAP. SPEC. VEI. =.164 4 L I T /GMU L E CO 1 HP RE f S S I S I L I TY =.722 PRESSURE = 1231 PSIA TFMPERATURE = 220 I)EG F ",ETHANE NOR-PFNTANF VAP. SPEC. VOL. =.3093 LIT/G.'OLE COMPRFSSIILITY =.837 VAPOR PH SE CU:'!PF.7400.2600 VAPOR PHAS f FJU: C]FF F 1.088,.1700 ~METhANE NOR-PENTANE LIOUID PHASE COMP.5320.4680 VAPOR PHASE CO;iP.8100.189s LIQUID PHASE F!J( CL EF 1.081 VAPOR PHASE FU(; rCLFF 1.(0001.3664 LIQUI ) ACT C' E- 1.(I6 1.032 LIQUID PHASE COMP.3060).6940 K ORS 1.351.56 LIQUID PHASE FUC CtEF.34.103 K CALC 1.543.493 LIQUID ACT C:]FF l. Igl.1I) I.08 ~L ~I~lJI 0D ACT CL~EF I.~ 0~91 1.~)!~ PERCENT I'! - 1().C,4 11.32 K OBS 2.647.273 K C4ALC 2.559.292 PERCENT DEV.3.34 -3.22

-122RUN NUMBER 27 PRESSURE = 1777 PSIA TEMPERATURE = 220 DEG F VAP, SPEC. VOL. =.1984 LIT/GMOLE COMPRESSIBILITY =.774 METHANE NOR-PENTAN E VAPOR PHASE COMP.7880.2120 VAPOR PHASE FUG COEF 1.0310.2314 LIQUID PHASE COMP.4560.5440 LIQUID PHASE FUG COEF 1.737.085 LIQUID ACT COEF 1.068 1.021 K OBS 1.728.390 K CALC 1.799.376 PERCENT DEV -4.12 3.47 RUN NUMBER 28 PRESSURE: 1501 PSIA TEMPERATURE = 220 DFG F VAP. SP-EC. VOL. =.2455 LIT/GMOLE C3OMPRESSIBILITY =.809 METHANE NOR-PENTANE VAPOR PHASF COMP.808(.1916 VAPOR PHASE FUG COEF 1.0094.2 993 LIQUID PHASE COMP.382().6180 LIQUID PHASE FUG CU(F 1. 86.092 LIQUID ACT COEF 1.079 1.013 K OBS 2.115.31( K CALC 2.123.313 PERCENT DEV -.37 -.81 RUN NUMBER 29 PRESSURE = 1260 PSIA TEMPERATURE = 220 DEG F VAP. SPEC. VOL. =.30(1 I T/GMtLE COMPRLSSIILI[TY =.834 fIF THANE F NOR-PENTANE VAPOR PHASE COMP. 8160.1845 VAPOR PHASE FUG COFF 1.0(?02.3640 LIQUID PHASE COMP.3100.6f;90) LIQUID PHASE FUG CiEF 2.299. 101 LIQUID ACT COEF 1.090 1.008 K OBS 2.63?.267 K CALC 2. ()0.280 PERCENT DEV 5'.01 -4.90 RUN NUMBER 30 PRESSURE = 1005 PSIA TEMPERATIURE = 220 DFG F VAP. SPEC. VOL. =.3894 LI T/GMOLE C[MPRESSIiL ITY =.859 PA THANE NOCR-PENTANF VAPOR PHASE COMP.8140.1863 VAPOR PHASE FUG COUF 1.000e,.4411 LIQUID PHASE COMP.2480.7520 LIQUID PHASE FUG COEF 2.807,116 LIQUID ACT COEF 1.099 1,005 K OBS 3.22.248 K CALC 3.083.264 PERCENT DEV 6.06 -6.63 AVE. ABS. PERCENT DEV. FOR SYSTEM = 3.963

METHANE ISOPENTANE BINARY SYSTEM RUN NUMBER 34 PRESSURE = 1899 PSIA TEMPERATURE = 220 DEG F RUN NUMBER 31 VAP. SPEC. VOL. =.1628 LIT/GMOLE COMPRESSIBILITY =.679 PRESSURE = 1256 P.SIA TFMPERATUKE = 220 DEG F METHtANE ISOPENTANE VAP. SPEC. VOL. =.2968 LIT/GMOLE COMPRESSIBILITY.819 VAPOR PHASE COMP.6860.3140 VAPOR PHASE FOG COFF 1.1471.1741 METHANE ISOPFNTANE LIQUID PHASE COMP.5660,4340 VAPOR PHASE COMP.7880.2120 LIQUID PHASE FUG COEF 1.652.096 VAPOR PHASE FUG COEF 1.0129.3h73 LIOUID ACT COEF 1.032 1.024 LIQUID PHASE COMP.3310.6690 K OBS 1.212.724 LIQUID PHASE FUG COEF 2.306.117 K CALC 1.487.563 LIQUID ACT COEF 1.055 1.006 PERCENT DEV -22.66 22.20 K OBS 2.381.317 K CALC 2.401.320 PERCENT DEV -.86 -1.07 RUN NUMBLR 35 PRESSURE = 1001 PSIA TEMPERATURE = 220 DEG F RUN NUMBER 32 VAP. SPEC. VOL. =. 38t,6 LIT/GMULE CIIMPRESSIRILITY =.850 PRESSURE = 1503 PSIA TEMPEFATIJRE = 220 DFG F METHANE ISOPENTANE VAP. SPEC. VOL. =.2376 LIT/GM[OLF C[)MPRt-SSI'I LITY =.784 VAPOR PHASF COMP.7910.2090 I VAPOR PHASF FIJG CFF) 1.()61.4495 [ METHANE I S) PL NT AN t. LI l)ID PHtASE C)MP.2620.7380 VAPOR PHASE COMP.174().028,0 LIUUI) PHIASE I-tJG CLE-F 2.817.134 VAPOR PHASE FUG CJEtF 1.0287.2954 LIQUID ACT C()FF 1.061 1.003 LIQUID PHASF CO'MP.3'60.60,sO K [BS.019.23 LIQUID PHASE FIJ, CClEF 1.9H3 K CALC 2.971.299 LIQUID ACT COEF!.(0'9.()() PERCENT UEV 1.59 -5.62 K OBS 1.',5.374 K CALC 2.()2.'63 PERCENT DEV -3.47 2.')) RUN NUMBER 36 PRESSURE = 75(' PSIA TEMPFRATURE = 220 DEG F RUN NUMBER 33 VAP. SPEC. VOlL. =.5?4' LIT/GMfLE COMPRFSSIILITY = 875 PRESSURE = 1721 PSIA TEMPt ATURE = 220 DEG F.ANE I SPNTANE VAP. SPEC. V{)l. =.1',fLI IT/(;G,"OLF CO)MPRFSSIBlEITY =.743 VAP)R PHASE CCMP.7 1,-,().7350 VAPOR PHASE FUi; CiFF 1.0037.5218 METHANE IPSlPF7NTANE LIQUIID PHASE C1OMP.191T.8080 VAPOR PHASE COMP.7460.2543) LIQUID PIHASE FUG CUFF h.639.162 VAPOR PHASE FiJ C(FF 1.00,4.2338 LIQUID ACT CUF 1.067 1.002 LIQUID PHASE C(M'P.454().546f0 K OHS 3.991.291 LIQUID PHASE FtJ~; CULF 1.780.10() K CALC.310 LIQUID ACT C)FF 1.043. PERCENT )EV.4 -.310 K OBS 1.641.405 PERCENT D)EV' ~.()4 -6'.6.3 KOBS 0. h4t3.465 K CALC 1. 7481.432 PERCENT DEV -6,.40() 7.06

RUN NUMBER 37 RUN NUMBER 40 PRESSURE = 499 PSIA TEMPERATURE = 220 DEG F PRESSURE = 1001 PSIA TEMPERATURE= 160 DEG F VAP. SPEC. VOL. =.8098 LIT/GMOLE COMPRESSIBILITY =.887 VAP. SPEC. VOL. =.3603 LIT/GMOLE COMPRESSIBILITY =.869 METHANE ISOPFNTANF METHANEE VAPOR PHASF COMP.7100.2900 VAPOR PHASE COMP.885M 9.1146 VAPOR PHASE FUG COEF 1.0205.6214 VAPOR PHASE FUG COEF56.4161 LIQUID PHASE COMP.1181.8820 LIQUID PHASE CO[MP.2830.7170 LIQUID PHASE FOUG COEF 5.453.22? LIQUID PHASE FUG C(EF 2.601.071 LIQUID ACT COEF 1.073 1.001 LIQUID ACT COEF 1.065 1.004 K OBS 6.012.329 K OBS 3.127.160 K CALC 5.735.35,q K CALC 2.905.172 PERCENT DEV 4.6-1 -9.95 PERCENT DEV 7.11 -7.42 RUN NUMBER 38 RUN NUMBFR 41 PRESSURE = 502 PSIA TEMPFRATURE = 160 DtG F PRESSURE = 1253 PSIA TFMPERATURE =160 DEG F VAP. SPEC. VOL. =.75't 1 IT/GMOLL COMPRESSIBILITY =.912 VAP. SPEC. VtL. =.277() LIT/GOLE COMPRESSIBILITY =.836 1tvtFT H-1 AN (\jE I SU3 PF F\1NTAN 1\1 E W~ETHANF ISOPENTANE VAPOR PHASE COMP.4'10.1593 VAPOR PHASF COMP.8790.1206 VAPOR PHASE FUG COFF.9'I 1.6045 VAPOR PHASE FG COFF.949).3291 LIQUID PHASE Ci[OlP.141,.8580 LIQUID PHASE COMP.3510.6490 LIQUID PHASE F(; CFf 4.933.115 LIQUID PHASE FG CO 2.147.06 LIQUID ACT COEF 1.078 1.001 LIQUID ACT COEF 1.058 1.007 K 0BS 5.931.186 K OBS 2.504.186 K CALC 5.42.3.190 K CALC 2.395.192 PERCENT DEV 3.'6 -2.39 PERCENT DEV 4.38 -3.18 RUN NUMBER 39 RUN NUMBER 42 PRESSURE = 755 PSIA TFMPFRATtJRE = 160 DEG F PRESSURE = 1505 PSIA TEOPFRATUR= 160 DEG F VAP. SPEC. VOL. =.48!(.) IIT/GlOLE- COMPRESSIBILITY =.887 VAP. SPEC. VOL. =.222( LIT/G(LE COMPRESSIBILITY =.805 XiETHABE I SDPENTANE MTHAN SENTANE VAPOR PHaSF CO!iP.8720).1283 VAPOR PHASE C!OIP.8690.1308 VAPOR PHASE FU;; CiE)-.8EF.82.001 VAPOR PHASE FUO C,)F F9467.2566 LIQUID PHASE C-:3MP.2180.7820 LIQUID PHASE COMP.418().5820 LIQUID PHASE FUG CCEF- 3.353.085 LIQUID PHASE FUG C.'EF 1.a83.057 LIQUID ACT CqEF 1.071 1.002 LIQUID ACT CUFF 1.051 1.012 K OBS 4. 0(0.164 K OBS 2.079. 225 K CALC 3.710.171 K CALC 2.058.225 PERCENT DEV 7.24 -4.13 PERCENT iEV 1.00 -.20

RUN NUMBER 43A- RUN NUMBER 46 PRESSURE = 1759 PSIA TEMP!RATiJRE = 160 )iL(; F PRESSURE = 511 PSIA TEMPERATURE 280 DEG F VAP. SPEC. VOL. =.1822 L IT/GOLE COMPRESSIBILITY.772 VAP. SPEC. VOL. =.8000 LIT/GMOLE COMPRESSIBILITY =.825 AFTHANE ISOPENTANE METHANE ISOPFNTANE VAPOR PHASE CiOMP.853').1474 VAPOR PHASE COMP.5200.4800 V POR PHASE FJG- C-)EF.9503.1953 VAPOR PHASE FUG COEF 1.0993.6221 LIQUID PHASE COiMP.489(.5110 LIQUID PHASE COMP.0916.9080 LIQUID PHASE FUO, CCEF 1.648.053 LIQUID PHASE FOlG COEF 5.619.364 LIQUID ACT COEF 1,044 1.018 LIQUID ACT COEF 1.069 1.000 K OBS 1.744.288 K OBS 5.677.529 K CALC 1.809.277 K CALC 5.462.585 PERCENT DEV -3.70 3.85 PERCENT,)EV 3.78 -10.72 RUN NUMBER 44 RUN NUMBER 47 PRESSURE = 1992 PSIA TFMP!:RATIJRF = 160'-)rF F PRESSURE = 759 PSIA TEMPFRATURE 280 nEG F VAP. SPEC. VOL. 1 CO`PRFSSIOILITY.728 VAP. SPEC. VOL. =.531517 LIT/GLEL COMPRESSIBILITY =.812,A F T H \AW I SJ) )PL F ATA NE FNI ET.HAN IE ISOPENTANE VAPOR PHASE COMP.R1~.1791 VAPOR PHASE COiBP.6030.397r) 9 VAPOR PHASE FlIG;,lEF.9756,.14?26 VAPOR PHASE FIJG C:JFF 1.0916.5290 LIQUID PHASE C(iM4P.544+).455 0 LIQUID PHASE COMP.1613.3390 LIQUID PHASE FULG COFF 1.509.()51 LIQUID PHASE FUG CO]tF 3.814.265 LIQUID ACT COEF I.n lq 1.024 LIQUID ACT COEF 1.064 1.001 K as 1. 5)6.394 K IOBS 3.738.473 K CALC L. 6,5.363 K CALC 3.71.502 PERCENT I)EV -.5 7.75 PERCENT DEV.56 -6.09 RUN NUMBER 45 RUN NUMBERE 48 PRESSURE = 2191 PSIA VFFlPlI-ATLjRF = 1.6, [)t F PRESSURE = 1001 PSIA TEMPERKTURE = 23 OefG F VAP. SPEC. VOL. =.1237 LtIT/GOLE C]tPRESSI;IIITY =.653 VAP. SPEC. VOL. =.38'9 LIT/GtJLE COMAPRESSIBILITY =.787 "M ITHA!iF I SOPFNTANE: tHAE I MFTHAAEH A iJ E I SOPFNTANE VAPOR P. E C0U'P.741).259)) VAPOR PHASE COOlP.6C360.36+40 VAPOR PHASE FUG CUEF 1.0655.0893 VAPOR PHASE FUG CiJF~ 1.1036.4519 LIQUID PHASF COtMP.633:).36,70 LIQUID PHASE CliiP).2310.7590 LIQUID PHASE FUG COEF 1.416.()4 LIQUID PHASE FUC C[EF?293h.217 LIQUID ACT COFF 1.028 1.038 LIQUID ACT COFF 1.058 1.002 K OBS 1.171.706 K OBS 2.753.473 K CALC 1.367.563 K CALC 2.815.482 PERCENT OEV -]15.77 20.22 PERCENT iEV -2.26 -1.74

RUN NUMBER 49 METHANE NEOPENTANE BINARY SYSTEM PRESSURE = 1267 PSIA TEMPERATURE = 280 DEG F VAP. SPEC. VOL..3031 LIT/GMOLE COMPRESSI!3ILITY =.775 RUN NUMBER 71 METHANE ISOPENTANE PRESSURE = 511 PSIA TEMPERATURE = 160 DEG F VAPOR PHASE COMP. 6510.3490 VAP. SPEC. VOL. =.7159 LIT/GMOLE COMPRESSIBILITY =.881 VAPOR PHASE FUG COFF 1.1044.3797 LIQUID PHASE COMP.3150.6850 METHANE NEOPENTANE LIQUID PHASE FUG COEF 2.370.186 VAPOR PHASE COMP.7610.2390 LIQUID ACT COEF 1.052 1.005 VAPOR PHASE FUG COEF.9991.5958 K OBS 2.067.509 LIQUID PHASE COMP.1532.8470 K CALC 2.257.493 LIQUID PHASE FUG COEF 4.849.174 PERCENT DEV -9.19 3.21 LIQUID ACT COEF 1.022 1.000 K OBS 4.967.282 K CALC 4.958.292 RUN NUMBER 49A PERCENT DEV.18 -3.47 PRESSURE = 1277 PSIA TEMPERATURE = 280,DEG F VAP. SPEC. VOL. =.2928 LIT/GH1OLE COiMPRESSIRILITY =.755 RUN NUMBER 72 METHA!JF ISOPENTANE PRESSURE = 763 PSIA TEMPERATURE = 160 DEG F VAPOR PHASE COMP.6439.357.) VAP. SPEC. VOL. =.4625 LIT/GMOLE COMPRESSIBILITY =.850 VAPOR PHASE FUG COEF 1.1301.3718F LIQUID PHASE C.0MP.300 6730) METHANE NEOPENTANE LIQUID PHASE FUG CGEF 2.354.185 VAPOR PHASE COMP.7970.2030 LIQUID ACT COEF 1.050 1. 005 VAPOR PHASE FUG COEF.9q883.4866 K OBS 1.948.533 LIQUID PHASE COMP.2320.7680 K CALC 2.187.501 LIQUID PHASE FUG COEF 3.321.129 PERCENT 0EV -12.27 5.89 LIQUID ACT COEF 1.020 1.001 K OBS 3.435.264 K CALC 3.426.266 RUN NUMBER 50 PERCENT DEV.28 -.76 PRESSURE = 1517 PSIA TEMPERATURE = 280 DEG F VAP. SPEC. VCL. =.2205 LIT/GMOLE C)OMPRESSIBILITY =.675 RUN NUMBER 73 MF-ETHA', E I SOPENTANE PRESSURE = 1005 PSIA TEMPERATURE = 160 DEG F VAPOR PHASE COMP.5810.4190 VAP. SPEC. VOL. =.34~1 LIT/GMOLE COMPRESSIBILITY = 828 VAPOR PHASE FUG COFF 1.24/+30.2779 LIQUID PHASE COMP. 438',8).5120 MFTHANF NEOPENTANE LIQUID PHASE FUG C2FEF 2.027. 158 VAPOR PHASE COMP 8190.1809 LIQUID ACT,,COFF 1.037 1.015 VAPOR PHASE FG CUOFF.9762.4076 K 01BS 1.191.I18P LIQUID PHASE CyOMP.3120.6880 K CALC 1.(91.614 LIQUID PHASE F(G CUEF 2.592.1(8 PERCENT DEV -42.03 25.03 LIQUID ACT COEF 1.010 1.002 K OBS 2.625.263 AVE. ABS. PERCt:NT' DLV. FtOR SYSTE = 1.719 K CALC 2.701.266 PERCENT DEV -2.91 -1.26

RUN NUMBER 74 RUN NUMBER 76B PRESSURE -1273 PSIA TEMPERATURE = 160 DEG F PRESSURE = 1709 PSIA TEMPERATURE = 160 DEG F VAP. SPEC. VOL. =.2573 LIT/GMOLE COMPRESSIBILITY =.789 VAP. SPEC. VOL. =.1617 LIT/GMOLE COMPRESSIBILITY =.666 METHANE NEOPENTANE METHANE NEOPENTANE VAPOR PHASE COMP.8130.1867 VAPOR PHASE COMP.7270.2730 VAPOR PHASE FUG COFF.9764.3189 VAPOR PHASE FUG COFF 1.0595.1707 LIQUID PHASE COMP.3910.6090 LIQUID PHASE COMP.5600.4400 LIQUID PHASE FUG COEF 2.119.095 LIQUID PHASE FOG COEF 1.683.082 LIQUID ACT COEF 1.015 1.003 LIQUID ACT COEF 1.011 1.007 K OBS 2.079.307 K OBS 1.298.620 K CALC 2.204.2?98 K CALC 1.605.486 PERCENT DEV -6.00 2.83 PERCENT DEV -23.64 21.69 RUN NUMBER 74A RUN NUMBER 77 PRESSURE = 1281 PSIA TEMPERATURE = 160 DEG F PRESSURE = 1748 PSTA TEMPERATURE = 160 DEG F VAP. SPEC. VOL. =.2552 LIT/GMOLE COMPRESSIBILITY =.787 VAP. SPEC. VOL. =.14Y4 LIT/GMULE COMPRESSIBILITY 25 METHANE NEOPFNTANE METHANE NEOPENTANE VAPOR PHASE COMP.8130.1874 VAPOR PHASE COMP.6850.3150 VAPOR PHASE FUG COFF.9766.3159 VAPOR PHASE FOG CliEF 1.1165.1458 LIQUID PHASE COMP.3970.6020 LIQUID PHASE COMP.6030.3970 LIQUID PHASE FUG COEF 2.109.094 LIQUID PHASE F!JG CFEF 1.655.082 LIQUID ACT COEF 1.015 1.003 LIQUID ACT COEF 1.009 1.009 K OBS 2.048.311 K nBS 1.136.793 K CALC 2.192.300 K CALC 1.49,.564 PERCENT DEV -7.04 3.69 PERCENT DFV -31.70 28.86 RUN NUMBER 75A RUN NUMBER 82 PRESSURE = 1521 PSIA TF.lMPERATIIRE = 160) DFG F PRESSURE = 310 PSI^ TE[!,PFRATURF- = 160 DEG F VAP. SPEC. VOL. =.200) LIT/GlOLE COMPRFSSIRILITY =.736 VAP. SPE(C. VOL. = 1.252 IIT/Gm CMPRFSSIIIITY =.900 -E Cl MPRFSSItlIt ITY =.0 M1F T H ANEN.PENANE TAE NE1PENTN' TANE VAPOR PHASE COMP. 7840).2160 VAPOR PHASF CE4-P ( 670 333 VAPOR PHASE FUG COFF.9968.?374 VAPOR PHASE FIJA, CJF 1.0140.6914 LIQUID PHASE COMP.4820.5180o LIQU1 0 PilASE {.n952p. 55 LIQUID PHASE FL-IG C(L:F 1.8H8.O87 LIQ UI ) PHAS Ft.(' CC! -F 1.916.262 LIQUID ACT COF 1.013 1.O05 LIQUID AC T COLFF 1.9 1.230 K OBS 1.627.417 KOBS 7.8 R'. 3t4 K CALC 1.8 H6.3 67 K CALC 7. 9,'7.38(0 PERCENT [)EV -14.83 11.98 PERCENT:JEV -( 2 -4. 31

RUN NUMBER 83 RUN NUMBER 86 PRESSURE = 308 PSIA TFIMPFRATURE = 220 DEG F PRESSURE = 1008 PSIA TEMPERATURE = 220 DEG F VAP. SPEC. VOL. = 1.2104 LIT/G:OLE COMPRESSIBIIITY =.819 VAP. SPEC. VOL. =.3501 LIT/GMOLE COMPRESSIBILITY =,775 MFT'THANE NEJPENTANE METHANE NEOPENTANE VAPOR PHASE COMP.3950.6050 VAPOR PHASE COMP.6700.3300 VAPOR PHASE FUG C(EF 1.1112.7007 VAPOR PHASE FIJG COEF 1.0718.4210 LIQUID PHASE C;MRP.0505.9490 LIQUID PHIASE CCMP.2820.7180 LIQUID PHASE FUG CCEH- 8.8)6.478 LIQUID PHASE FUG CCEF 2.7'-)99.187 LIQUID ACT COEF 1.02 1.100 LIQUID ACT ClEF 1.017 1.001 K OBS 7.F'22.638 K UBS 2.376.460 K CALC 8.047.682 K CALC 2.656.445 PERCENT DEV -3.52 -6.90 PERCENT )EV -11.78 3.25 RUN NUMBER 84 RUN NUMBER 8'7 PRESSURE = 50)3 PSIA - TEMP:FRATURE = 220 DEG F PRESSURE = 1251 TSIA TFMPERATURE = 220 DEG F VAP. SPEC. VOL. =.7468 LIT/GMBOLE COMPRFSSIBILITY =.825 VAP. SPEC. VOL. =.2627 LIT/GM!OLE CJMPRESSIBILITY =.722 MFTiHANF NEnPENTANE MFTHANE NEOPENTANE VAPOR PHASE COMP.5f-30.4370 VAPOR PHASE CORP.t540.3460 VAPOR PHASE FUG COEF 1.0752.6055 VAPOR PHASE FUG COEF 1.1122.3341 LIQUID PHASE COMP.116,1.'8.830 LIQUID PHASE COMP.3770.6230 LIQUID PHASE FUG COEF 5.410.-14 LIQUID PHASE FUG C(EF 2.314.164 LIQUID ACT COEF 1.020 1.000 LIQUID ACT COEF 1.014 1.002 K OBS 4.820.495 K 08S 1.735.555 K CALC 5.135.518 K CALC 2.110.491 PERCENT DEV -6.52 -4. 73 PERCENT l)EV -21.64 11.59 RUN NUMBER 85 RUN'NUMBER 88A PRESSURE = 748 PSIA TE:PPFATtIRE = 220 D[E F PRESSURE = 1434 PSIA TFMPERATURE = 220 DEG F VAP. SPEC. VL. =.547'3 LIT/GMICLF CWi'4PRESSBIL ITY =.900 VAP. SPEC. VOL. =.200)4 LIT/Gi4OLE CliMPRESSIBILITY =.631 i,1F T I ANEF N J U P F PN T ANE' E TH ANE EP i'PENTA NE VAPOR PHASE COMP,:9!). h610 VAPOR PHASE CCMP.5850).4150 VAPOR PHASE FlJ, C'JEF.~ 655.4934 VAPOR PHASE FIJG CDEF 1.2384.2498 LIQUID PHASF C;r'P. 1 97L. Q9(03 LIQUID PHASE COMP.4710).290 LIQUID PHASE FOIG ItiU t.29!) LIQUID PHASE F(, COEF?.0, 2.152 LIQUID ACT Cf)EF 1.(1I9 1.000J LIQUIDO ACT COFF 1.01 1.004 K OBS...24 /. ~50 K OBS 1.242.7,4 K CALC.467 K CALC 1. 8d5.610 PERCENT DEV -20.C0; -3.;4 PERCENT DtV -3D'.6 22.24

RUN NUMBER 91 METHANE NEOPENTANE NORMAL PENTANE TERNARY SYSTEM PRESSURE = 506 PSIA TEMPERATURE 280 DEG F VAP. SPEC. VOL. =.6733 LIT/GMOLE COMPRESSIBILITY =.687 RUN NUMBER 95 METHANE NE]PENIANE PRESSURE = 503 PSIA TEMPERATURE = 160 DFG F VAPOR PHASE COMP.2800.7200 VAP. SPEC. VOL. =.7634 LIT/GMULE COMPRESSIBILITY =.925 VAPOR PHASE FUG COEF 1.2835.6102 LIQUID PHASE COMP.0683.9320 METHANE NEOPENTANE NOR-PENTANE LIQUID PHASE FUG COFF 5.674.487 VAPOR PHASF COP.8450.0577.0972 LIQUID ACT COEF 1.020 1.000 VAPOR PHASE FUG COE.9792.6428.5902 K OBS 4.100.773 LIQUID PHASE CO1MP.1407.2160.6430 K CALC 4.508.798 LIQUID PHASE FUG COEF 4.924.176.095 PERCENT 0EV -9.96 -3.33 LIQUID ACT CiEF 1.089 1.044 1.013 K OBS 6.006.267.151 K CALC 5.478.286.163 RUN NUMBER 92 PERCENT DEV 8.79 -7. 09 -7.87 PRESSURE = 755 PSIA TEMPERATURE = 280 DEG F VAP. SPEC. VOL. =.4471 LIT/GMOC)LE COMPRESSIBILITY =.681 RUN NUMBHFR'6 EETHANF NEOPENTANE PRESSURE = 7R1 PSIA TFPERfATUJRE = 160 DEG F VAPOR PHASE COMP.4070.5930 VAP. SPEC. VOL. =.4975 LIT/SMJLE COMPRESSIBILITY =.900 VAPOR PHASE FUvG COEF 1.26.1.5063 LIQUID PHASE COMP.1632. 8370 MF THAQE NFEPENTANE NOR-PENTANE LIQUID PHASE FOG CO!EF- 3.834 349 VAPOR PHASF CO8P.71).0459.0827 LIQUID ACT COEF 1.018 1.000 VAPOR PHASE FtO] f-.;LtF.970q.5443 4830 K OBS 2.494.708 LIQUID PHASE C(oMP.2f60.1997.594 K CALC 3.077.690 LIQUID PHASE FOG C(,FF A.370.131.071 PERCENT DEV -23.q 2.60 LIQUID ACT COEF I.0IP 1.037 1.017 K OBS 4.22.230.139 K CALC 3. 757.250.149 RUN NUMfBER 930 PERCENT D)EV 11.15 -8.55 -7.22 PRESSURE = 1004 PSIA TFaPFRATUJRE = 280 DES F VAP. SPEC. VOL. =.30213 LIIT/GMOLE CVPPFSSIOII ITY =.612 RUN NUMB[fR 97 i'FTHANE NF LPFNTANE PRESSURE = 25I P SIA T:NP: kATURF = 160 DE G F VAPOR PHASE CUMP.416 O.5840 VAP. SPEC.. VO;L. =.,8 l0 I. T/GtOLE COMPRESSII1L ITY =.847 VAPOR PHASE FUG COFF 1.3795.4005 LIQUID PHASE COMP.28 1).7190 MTHA"NG NEOPENTANE NOR-PENTANE LIQUID PHASE FUG COEF 2.927.2 1 VAP,]R PHAS C;MP.78.04(4.0818.;i7f: JO.()40 4.081 LIQUID ACT COEF.015 1.001( VAPUJ, PHASE FUO J..375.3059i K 083BS 1.4%. 12 LIQUID PHASE CFIIF'. 7i).164().4980 K CALC 2. 15.70% LIQ(UI D PtiAFSE FOlG F l.096.052 PERCENT DEV -45.5 13.45 LQUID AT F..4 1.029 LIQUJID ACT C! _IE c;.C- 1.024 1 0 2 K UBS. ^ 5.246.164 AVE. ABS. PERCENT DEV. rF!P SvST-,i = 11.594 K CALC 2 i'.61.174 PERCENT DEV.1 -5.83 -6.20

RUN NUMBER 98 RUN NUMBER 101 PRESSURE = 1505 PSIA TEMPERATURE = 160 DEG F PRESSURE 2120 PSIA TEMPERATURE = 160 DIGF VAP. SPEC. VOL. =.2254 LIT/GMOLE COMPRESSIBILITY.817 yAP. SPEC. VOL. =.1316 LIT/GMOLE COMPRESSIBILITY =.672 METHANE NEOPENTANE NOR-PENTANE METHANE NEOPENTANE NOR-PENTANE VAPOR PHASE COMP.8700.0410.0891 VAPOR PHASE COMP.7750.0591.1655 VAPOR PHASE FUG COEF.9626.3039.2340 VAPOR PHASE FUG COEF 1.0804.1346.0760 LIQUID PHASE COMP.4000.1482.4520 LIQUID PHASE COMP.6020.0976.3000 LIQUID PHASE FOG COEF 1.853.087.047 LIQUID PHASE FUG COEF 1.447.076.041 LIQUID ACT COEF 1.061 1.018 1.036 LIQUID ACT COEF 1.036 1.002 1.076 K OBS 2.175.277.197 K OBS [.287.606.552 K CALC 2.043.292.209 K CALC 1.388.564.575 PERCENT DEV 6.09 -5.50 -6.07 PERCENT DEV -7.82 6.82 -4.16 RUN NUMBER 99 RUN NUMBER 102 PRESSURE = 1759 PSIA TEMPERATURE = 160 DIG F PRESSURE = 1006 PSIA TEMPERATUJRE 160 DEG F VAP. SPEC. VOL. =.1851 LIT/GMULE CUMPkFSSIBII.ITY.784 VAP. SPEC. VOL. =.3F11 LIT/GOLE- COMPRESSIBILITY =.875 METHANEF NEOPENTANE NOR-PENTANE IFTHANE NEOPINTANE NOR-PENTANF VAPOR PHASE COMP.8550.0429.1025 VAPOR PHASE COMP.8 790.0406.0807 H VAPOR PHASE FUG OlF.9711. 2396.1717 VAPOR PHASE FOG CUFF.9626.4522.3854 LIQUID PHASE COMP.4610.1314.4080 LIQUID PHASF COMPB.2780.17(0.5460 LIQUID PHASE FOG COEF 1.648.(Ri.044 LIQUID PHASE FU(0, C:EF 2.569.108.059 LIQUID ACT COEF 1.054 1.01? 1.045 LIQUID ACT CUFF 1.075 1.030 1.023 K OBS 1.655.326.251 K OBS 3.16?.231.148 K CALC 1.766.3444.267 K CALL 2.693.247.156 PERCENT DEV 3.60 -5. 29 -6.41 PERCENT D.FV -6.95 -5.42 AVE. ABS. PEkCINT I0EV. FOr SYSTEM = 6.382 RUN NUMBER 100 PRESSURE = 2013 PSI1 TIMPI PATLR[ = 160 0)F0 F VAP. SPEC". VO)L..1463 LIT/GU'Ll. CWAPRFSS 161 LlY =.I70' MF- THAYF Nf I)PFNT A8JF NOR-PENTANF VAPOR PHASF CGUMP. (.6306.1413 VAPORI PHASE FUG CUFF IGII'1.6.1008 LIQUID PHASE LU-MP 50".1112 333<;0 LIQUID PHASE 0I.J(0 CCVIFI I 4'-' S.077.041 LIQUID ACT CGI-E I. 105 1.063 K 085 1. 464.4i32.417 K CALC 1 614.476.437 PERCENT DFV 1..41 -4.90

METHANE ISOPENTANE NO'RMAL PL,',T,;NF TEPNJARY SYSTEM RUN NUMBER 54 RUN NUMBER 51 PRESSURE = 1493 PSIA TFMPERATURE = 160 DEG F PRESSURE = 504 PSIA TFMPLPRATURF = 160 DEG F VAP. SPEC. VOL. =.2294 LIT/G.OLE COMPRESSIBILITY =.825 VAP. SPEC. VOL. =.7619 LIT/GCll)LE COMPRESSIBILITY =.925 METHANE ISOPFNTANE NOR-PENTANE METHANE ISOPFNTANE NOR-PENTANE VAPOR PHASE COMP.8890.0307.0803 VAPOR PHASE COMP.8.710.03<95.0892 VAPOR PHASE FUG COEF.9365.2765.2598 VAPOR PHASE F(JG COFF.9c15:.6244.6102 LIQUID PHASE COMP.4060.1536.4410 LIQUID PHASE COMP.13862.2230.6390 LIQUID PHASE FUG COEF 1.865.057.047 LIQUID PHASE FUC CCEF I. 14.115.095 LIQUID ACT COEF 1.075 1.002 1.023 LIQUID ACT CUFF 1.113 1.002 1.004 K OBS 2.190.200.182 K OBS 6.2.84.177.140 K CALC 2.140.208.187 K CALC 5. 6J8.184.156 PERCENT DEV 2.27 -4.02 -2.50 PERCENT DEV 10.7' -3.81 -11.78 RUN NUMBER 55 RUN NUMBER 52 PRESSURE = 1975 PSIA TEMPERATURE = 1'60) DEG F PRESSURE = 755 PSIA TFMPFkATURE = 160 DE:, F VAP. SPEC. VOL. =.1590 LIT/GYOLE COMPRESSIBILITY.756 VAP. SPEC. VOL. =.4949 LIT/GMOLE COMPRESSIBHILITY =.900 METHANE ISOPENTANE NOR-PENTANE H METHANE IS]SOPENTANE NOP-PENTANE VAPOR PHASE CUMP.8490.0405.1107 VAPOR PHASE COMP.894O.1)3")5.0755 VAPOR PHASE FUG COE.9529.1600.1452 VAPOR PHASE FUG COE9-2.9219.5190.37 LIQUID PHASE COMP.504f).1290.3670 LIQUID PHASE CO)RP.2110.2060.5830 LIQUID PHASE FUG C[EF 1.518.051.042 LIQUID PHASE FUG CCFF 1.353.085.071 LIQUID ACT COEF 1.060o 1.007 1.038 LIQUID ACT COEF 1.103 1.:) 01 1.007 K OBfS 1.485.314.302 K OBS 4.2 )7.14:.13o) K CALC 1.68..320.299 K CALC 3.840.164.141 PERCENT DEV -.20 -1.80.97 PERCENT DEV'?. 7 -10.98 -9.02 RUN NUMBER 55A RUN NUMBER 53 PRESSURE = 1995 PSIA E-PERATURE = 14)0 DEG F PRESSURE = 10')3 PSIA TFMPrFATTJRE = 160 OEG F VAP. SPEC. VOL. =.1551 _IT/GOi)LE COMPRESSIBILITY =.745 VAP. SPEC. VOL. =.36'? L I/T;L.'ALE CUOPIR!SSIILITY =.875 AETHANE I SOPENTANE, NOR-PENTANE ".A1 T1A' I SidPf.TANE NORl-PEfNTANE VAC,) PHASE CO.,P.8420.0422.1159 VAPOR PHASE CORMP.12(,"I I.8 ).'46.0733 VAPa]R PHASF FLGJ C']EF.9618.1523.1377 VAPOR PHASE FU(, Cl)[F.O 12.4'240.4072 LIQoJli) PHASE CO.RP.5210.1229.3560 LIQUID PtHASE CM8P.2 74'.1I 0;8.537(0 Li!1D PHASE F!JP CrIEF 1.507.051.042 LIQUID PHASE FPJ CUG'r-?.0[.:71.059 LI(lJ[:l) ACT C(]F 1.057 1.008 1.041 LIQUID ACT COEF 1.(4 1. 0Y 1.011 K [ORS 1.616.343.326 K OBS'.,74.157.136 K CALC 1.6b17.335.315 K CALC 2.e,.16.146 PERCENt'EV -2.51 2.48 3.35 DERCENT DEV -..) -F.V -7.01

RUN NUMBER 58 RUN NUMBER 61 PRESSURE = 2208 PSI A TEMPERATURE = 150 DEG F PRESSURE = 1519 PSIA TEMPERATURE: 220 DIG F VAP. SPEC. VOL. =.1214 L!T/GMOLE COMPRESSIBILITY =.663 VAP. SPEC. VOL. =.2407 LIT/GMOLE COMPRESSIBILITY =.803 METHANE I SOPENTANE NOR-PENTANE METHANE I SUPENTANE NOR-PENTANE VAPOR PHASE COMP.75~0.0632.1783 VAPOR PHASE COMP.8010.0527.1466 VAPOR PHASE FUG COIF 1.0547.OB90.0772 VAPOR PHASE FUG COEF 1.0132.3091.2907 LIQUID PHASE COMP.5930.1039.3030 LIQUID PHASE COMP.B890.1540.457Q.... LIQUID PHASE FUG COIF 1.385.048.040 LIQUID PHASE FUG COIF 1.~66.106.Og2 LIQUID ACT COIF 1.046 1.016 1.057 LIQUID ACT C[IEF 1.070 1.001 1.019 K OBS 1.278.~08.588 K OBS 2.059. 342.321 K CALC 1.374.549.541 K CALC 2.077.343.322 PERCENT I)EV -7.4.~, g.77 8.08 PERCENT DEV -.8,9 -.21 -.30 RUN NUMBER 59 RUN NUMBER 62 PRESSURE = 1765 PSIA 1E~'~P.F-RATURE = ~2() OFG F PRESSURE = 1263 PSIA TEMPERATURE = 220 DEG F VAPo SPEC. VOL. =.lg57 L[T/G'tCLE CL)MPRESSI[I[LITY =.75g VAP. SPEC. VOL.:.2997 LIT/GMOLE COMPRESSIBILITY:.831....'qF TvtANE I q.qP[-NTANE I',,r)~- PFNTANE METHANE [ SOPENTANE NOR-PENTANE VAPOR PHASE COHP.7719.{)(1l~.1661 VAPOR PHASE COMP.8100.0509.139l I VAPOR PHASE FLJG COEF 1.0467.7%91.2212 VAPOq PHASE FUG Cf. JEF 1.0044.3791.3607 LIQUID PiJASE C~]MP.4~4;:). 1391.4070 LI~IJID PHASE CUMP.3190.1716.5090 k,.N LIQUID PHASE FUG CC, EF 1.746.,')gO.,'18(~ LIQUIIg PHASE FUG COEF 2.295.117.101 ro LIQUID ACT COEF. 1.061 1.003 1.027 LIQUID ACT COiF 1.080 1.000 1.013 I KOBS 1.09:9.441.413 K UBS 2.539.297.273 K CALC 1.171.414.397 K CALC 2.467.30B.284 PERCENT DEV -4.~,} ~,.q9:~.~8 PERCENT L)EV 2.86 -3.77 — 3.97 RUN NUMBER 50 RUN NUMBER 63 PRESSURE = 2047 PSIA TE'.4P-.:~ATURF = 2'7() OFG F PRESSURE = 995 PSIA T[:,"4. PERATURE = 220 DEG F VAP. SPEC. VOL. =.1~,.~'~ I IT/:VJLE C~}~,~;)~,FSS[GItITY =.733 VAP. SPEC. VOL. =.393~ LtT/G:~OLE COMPRESSIBILITY =.85g ~,~r T~tA>~F I SJPF NlANF NC]i~-PFNTANE METHANE I SUPFNTANE NOR-PENTANE VAPOR PHASE C[)~qP.747~).0652.1P78 VAPOR PHASE C(l~,,P.810().0510.1394 VAPOR PHASE FLJCi CCJFF 1.,t7':~.l~9,'+.1725 VAPOP, PHASE FU(; C!IFF 1.()(109.4613.4435 LI'~IJIF) PHASE CgMI).5~,-,.).! 117.3330 LIc~UID P~tASF (,(~MP.251t).187~.5610 LIQUID PHASE FIJO C(:FF l.~-(,3.o93.08() LI(JUID PHASE FUC, [L~ "F 2.P3~.135.117 LIQUID ACT COIF 1.',]"~ L.010 1.043 LIQUID ACT C1F[: 1.0~9 1.000 1.008 KOBS l.~t,,'-..~34.56q K gBS 3.227.272.248 K CALC I.c2 ~.~+gS./-,R~ K CALC 3.090,.292.265 PERCENT r)EV -1.~. 12 15.,~4 1 ~.79 PERCENT LJFV ~..55 -7.37 -6.81

RUN NUMBER 64 RUN NUlMBFR 67 PRESSURE = 753 PSIA TEAP..rRATURE= 220 OE(; F PRFSSURE 7,,,C PS 1A I'-,PEPATURF 280 DEG F VAP. SPEC. VOL. =L52)1/T/cACL-CIE CU.MPRF:SSV3-ILITY =.875 VAP. SPEC. V:-lJL. =.3'_3,8 L ITc/C'f/GMOL E COMPRESSIBILITY =.819 v. F T;,\?,F IlSOENTA:4E N..3R-P E NTANF MFTHANF I SOPENTANE NRPN VAPOR PHASF COF.'P. 7 i80.577. 154 5 VAPOR PHASE C(}MIP t. 0 23 VAPOR PHASE FOG Ci3[F I1.Cn~-3".53 75. 520 3 VAPOR PHASE FUG CLF-''F 1.();H 72 L IQU ID P HAS E CUMP.1 8 67. 20 5).6090 LIQUID PHASE C-JvPMI022.20206-3 LIQUID PHASE FUG CUF:F 3.666. ] 6 3.141 ~l,41JU.r) PH.ASE FUG CF. F'..()q). 265.3 LIQUID ACT COE-F I1.09~o7 1 I.0O 1005 LIQUID ACT CfIEF I1.092 1.002104 K OBS 4. 2 2 1 2 8 1 2 54 K O~~~~~K BS 3. 8~4. 479.4 29 K CALC 3.c:'f.30,__; 3.2 72 K CALC -3.F-, 25.495.463 PERCENT DEV 5.5f7 -7.AO -7. 32 PERCENT!.EV I 3. 5~)I7 RUN NUMBER 65 RUN NW~GE.*6t PRESSURE = 57 PS It TF"%,PFP.ATl.JF:._0!)[G- F RSUR= 0)'?' TF.~Pl'R.ATtRE = 2.0 F)E G F VAP. SPEC. VOL. =.R)7 L I T / Gv'.1 E C(]'MP K 7S S I iRI I ITY =.) 4 VA P. SP EC. V,.JL. L.I.( LT/,'.I:]LE C[]MPRFSSI:SILlTY =.800 AF' TI-A" I S A?- P.~.' NTANF N'G r - P[L~ NT A NF'q:': (-'-H AF' I SOPFNTANF NOR-PENTAN VAPOR PHASE CE.M'P. 7 i44~;.7()1..1')4 VAPO,<~ PHASEC [':C. -:2.1.;)R66h —1 VAPOR PHASE FU,.q C~F F 1.')j1 7 4 6f, ".1 64 VAPOq PHASe F U'L:iF1. ('SI LI~~u~n PHASE CfWP - I I I.1 ~~~~~~' ~~- 1. (Ki 1 7 ~.6 13.441 6 LIQUID~~~~~~~~~~~~~~ PtA 2~F.12'2. 22()0.6600 LI1'.JU I D,'-q~A'E C.:P LIQUID PHASE FtJC. Cf.FF 5. gr,.I2(.10 LI6J PiS FU(2 C~tF _. LIQUID ACT Ci[-:.IF I. 105 L.) I.0 LQLJ I: PACT C.[F i tFF17 9 I0 5 3 I -1) 0 3 ~~~~~~~~~~~~~~~~l.I('I CT, I 4.000 1.007 K OBS Il-,;:) P 1..8 K g0 2.,~6.464 43 K CALC 5.i4r. gqe. 09 K CALC 2.c;4.47l.446 PERCENT OEV 1.~ -).3 -1. 7 PERCENT!DFV 91- R UN N U WEN,5 RUN NUMBER l~ VPRS UE=:,e SA'.'-:Rf!F=2oDSF PRESSURF =1263i,SIA IF8PRALR = 2H OG F VAP. SPEC.. VOL.:.~(-42 l. IT/f;'~i)LE Cr)M~~~~~i~FSSIRILITY =.8'~[ VAP. SPEC. VGL.VAP =SP.C. -,,+ L I1I L-I }I' F CC1.PF C S I R[LITYIL=T.,800 T"'f: TtA""i-F I S'.]P! N T ANF N R~- P F NTANF:'I A {A",' I SJPENTANE NOR-PENTAN VAP'O', PHASF C(';'QP.I.''.11'2.")20 0 VAPO'~ PHASE. C(',:,'P.-'7 4,)30.2408( VAPOR PHASE FU'.~7 C;L}[F 1.;C';".:'. 6 17 1.99 90 VAPO3R PhASc_ I-LJG C'miI-F 1.;:, I.39,'4.370g9 LIQUID PH'ASE CO')t.'[-.0-,n ~.22r'0.61 AI0 L I.:JIJID P I`A-) 3ASE14452 C( F! F 1')69~1/ 4 5g LQU ID P HAS E Ft.J( C(.'[I-:. -. 3%I').4,.3 1, LIQUID PHASE FL(:. C"fF 2. 4,':~ LIQUID ACT CU.EF I. 0'.B 1..'0 03 1. 002 LIQUID ACT Ci';r_. " - ].()751.001 KOB8S'(-"4 1. "I,.47 0 K OBS 2.2).492.463 ~~~~~~~~~~~~~~~~~~~~~~~K CALC 2.524 CA.481.462 PERCENT f')EV 5. ('/.') -?7. 2 - 11. 44 PRETIE 3 02

RUN NUMBER 66A RUN NUMBER 69A. PRESSURE = 541 PSITA TEMPERATURE= 280 BEG F PRESSURE= 1255 PSIA TEMPERATURE 280 DEG F yAP. SPEC. VOL..7614 LIT/GMOLE COMPRESSIBILITY=.831 yAP. SPEC. VOL..3072 LIT/GMOLE COMPRESSIBILITY.778 METHANE ISOPENTANE NOR-PENTANE METHANE ISOPENTANE NOR-PENTN VAPOR PHASE COMP.5680.1197.3120 VAPOR PHASE COMP.6740.0870.238 VAPOR PHASE FUG COEF 1.0925.6187.6007 VAPOR PHASE FUG COEF 1.1046.3924.372 LIQUID PHASE COMP.1030..2320.6650 LIQUID PHASE COMP.3040.1797.516 LIQUID PHASE FOG COEF 5.310.347.31? LIQUID PHASE FOG COEF 2.390.188.169 LIQUID ACT COEF 1.097 1.003 1.002 LIQUID ACT COEF 1.075 1.000 1.011 K OBS 5. 51 5.516.469 K OBS 2.2 17.484.461 K CALC 5. 3 33.563.521 K CALC 2.3 25 1.478.459 PERCENT DEV 3.28 -9.13 -11.00 PERCENT [DEV -4. 87 1.8.39 RUN NUMBER 67A RUN NUMBER 70B PRESSURE 757 PS16 TEMPFPATURE = 2O DEG F PRESSURE.= 1565 PSIA TEMPI-RATURE 290 BEG F VAP. SPEC. VOL. *5359 LIT/MOMULE CoMPRESS191I.ITY=.819 VAP. SPEC. VOL..21H7 LIT/GMOLE COMPRESSIBILITY=.691 ME THIiANE ISOPFNTANE NOR-PE:NTANE METHANE ISOPENTANE NOR-PENTN VAPOR PHASE COMP. 6 290 to02(.2690 VAPOR PHASE COMP.6160.0994.285E VAPOR PHASE FOG COLE 1.0k874.593 77.51F87 VAPOR PHASE. FOC CoLC 1. 219 1.2822. 261 LIQUID PHASE COMP.1642.2140.6210 LIUD HS C NP.4530.1384.408CLIQUID PHASE FOG COEF 3.824.266).239 LIQUID PHASE FlUG COEF 1.975.165.149 LIQUID ACT COEF 1. C01? 1.001 1.00)4 LIQUID ACT COEF 1.05 I' 1.003 1.024 K 085 3.h6i1.477. 433 K OBS 1. 3 60.718.699 K CALC 3. 6 3 6.495.463 K CALC 1. 711.588.584 PERCENT 0EV -.13 -3'3 -6.99 PERCENT 0EV - 25. 84 18.15 1b.35 AVE. ABS. PERCENT DEV. FOR SYSTEM 6.073 RUN NUMBFR 68A PRESSURE 1031 PSIA ]TEMPERATU.REF?,R0() BEG F yAP. SPEC. VOL...3830' LIT/GlMOLE C UA P.k SSI6ILTY=.797 ME iHAUF I SJ Pf PFAN E ~- P F t..T NE VAPOR PHASE COMP.6650.000.440 VAPOR PHASE FUG CUFF 1.0936..45)30.4332 LIQUID PHASE ClUMP. 24 2'0.1%46.5630) LIQUID PHASE FIJG COLEF 2. p56.1I ~.1(2 LIQUID ACT COLE 1.082 1. 090 1.007 K OBS. 7 48.465. 433 K CALC 2.826.47').446, PERCENT DEV -2.64 -1.07-30

APPENDIX B EXPERIMENTAL DATA Table XXVI presents the experimental data for all three binary systems considered, in this research. Data are arranged such that the values presented with an "A" suffix are the averaged values of duplicate or triplicate analyses. Beneath the averaged values are listed the individual analyses of each run, Also included in this table are the equilibrium pressure and temperature for each run, Table XXVII presents the experimental data for the two ternary systems investigated. As described in the previous paragraph, averaged compositions values are denoted with an "A" suffix~ Pressures and temperatures are included for each run. -135

-136TABLE XXVI EXPERIMENTAL DATA FOR BINARY SYSTEMS

METHANE NORMAL PENTANE OLNAkY SYSTEM METHANE NORMAL PENTANE BINARY SYSTEM RUN PRESS TEMP VAPOR COMPOSITIUN LIQUID COMPOSITION RUN PRESS TE.MP VAPOR COMPOSITION LIQUID COMPOSITION NUMBER PSIA (F) CI N-C5 Cl N-C5 NUMBER PSIA (F) Cl N-C5 Cl N-C5 21 1502 220.8080 A.1920 A.3804 A.6196 A 26 1999 220.7400 A.2600 A.5316 A.4684 A.8011.1989.3804.6196.7440.2560.5323.4677.8096.190't.7360.2640.5309.4691.8135.1865 27 1177 220.7883 A.2117 A.4559 A.5441 A 22 1265 220.6115 A.1885 A.3242 A.6758 A.7835.2165.4573.5427.8061.1939.3221.6779.7909.2091.4545.5455.8169.1831.3264.6736.7906.2094 23 1231 220.8102 A.1898 A.3059 A.6941 A 28 1501 220.8084 A.1916 A.3825 A.6175 A.8066.1934.30173 t9217.8083. 1917.3829.6171.8211.1789.3053.6947.8086.1914.3928.6072.8028.197~ 3050.6950.371 7 b17 283 24 1023 220.8059 A.1941 A.2527 A.7473 A 29 12s0 220.8155 A.1845 A.3103 A.6897 A.7994.200o 22521.7479.8203.1797.3129.6871.8124.1876.2532.7468.8108.1892.3078.6922 25 1001 220.8047 A.1953 A.2472 A.7528 A 30 1005 220.8137 A.1663 A.2481 A.7519 A.8008.1992.2470.7530.'8123.1877.2493. 7507.8079.1921.2474.752o.8151.1649.2469.7531.8053.1941

METHANE ISO-PENTANE BINAKY SYSTEM METHANE ISU-PENTANE bINARY SYSTEM RUN PRESS TtMP VAPOR CLMPUSITIUN LIQUID CUMPOSITION RUN PRESS TEMP VAPOR COMPOSITION LIQUID COMPOSITION NUMBER PSIA (F) CI ISOC5 Cl ISOCS NUMBER PSIA IF) CI ISUC5 CI ISOC5 31 1256 220.7816 A.2124 A.3310 A.6690 A 36 159 220.1653 A.2347 A.1917 A.8083 A.7862.2138.3251.6749.7646.2354.1901.8099.7891.2109.3328.6672.7660.2340.1897.8103.3368.6632.1955.8045.3293.6707 37 499 220.7104 A.2896 A.1181 A.8819 A 32 s0 3 220.7742 A.2258 A.3900 A.6040 A.7120.2880.1185.8815.7746.2254.3946.6054.7082.2918.1185.8815.7138.2262 -3974.6026.7110.2890.1173.8827 33 179I 22u.7462 A.2538 A.4536 A.5464 A 38 502 160.6407 A.1593 A.1418 A.8582 A.7440.5o50.4500.5500.6390.1604.1411.8589.7484.2516.4568.5432.841I.1562.1416.8584.4539.5461.1429.8571 34 1899 22 0 u861 A.3139 A.5602 A.4338 A 39 155 IU.16717 A.1263 A.2182 A.7818 A.6811 31 6d.5oo2.4336.87727.1213.2198.7802.b911.3069..)6o3.4331.6108.129.2.2169.7831.66b1.31i9.2190.7810.2172.7828 35 1001 220.1907 A.2U93 A.2617 A.1383 A.7698. 21 02.2648 73352 40 1001 1o0.8854 A.1140 A.2829 A.7171 A 1(91lo.2084.2608.7392.8861.1133.2836.7164.z594.740b.8177. 113.2826.7174.2619.7361.8811.1103.2825.7175 41 12'3 10.8 794 A.120O A.3513 A.6487 A.6192.1208.3508.6492.di7o.1L'04.3534.6466.34 97.503

METHANE ISO-PENTANE BINARY SYSTEM METHANE ISO-PEONTAtNE BINARY SYSTELM RUN PRESS TEMP VAPOR COMPOSITION LIQUID COMPOSITION RUN PRESS TEMP VAPOR COMPOSITION LIQUID COMPOSITION NUMBER PSIA (F) C1 ISOCS C1 ISOCS NUMBER PSIA (F) C1 ISOC5 C1 ISOC5 42 1505 160.8692 A.1308 A.4178 A.5822 A 48 1001 280.6360 A.3640 A.2313 A.7687 A.8687.1313.4178.5822.635o.3644.2302.7698.8698.1302.4179.5821.6405.3595.2325.7675.6319.3681 43A 1759 loO.8520 A.1474t A.4889 A.5111 A.8477.1523.4899.5101 49 1267 28U.8510 A.3490 A.3152 A.6848 A.8554.1446.4879.5121.b463.3537.3151.6849.8546.14)4.b o50.3444.3152.6848 44 1992 lou.b209 A.1791 A.5451 A.4549 A 49A 1271 260.6432 A.3568 A.3301 A.6699 A. 8136.I8o4.5454.4546.6432.3568.3300.6700!.8253.1141.5448.4552.3302.6698.8231.1 o3 D 50 1517 280.5814 A.4186 A.4881 A.5119 A 45 191 1o0.7413 A.Z5b7 A.6333 A.36o7 A.5800.4200.4867.5133.7409.Z591.6328.3672.5o28.4172.4857.5143.7417. 2583.6361.3639.4918.5082.6311.3689 46 511 260.5204 A.4195 A.0916 A.9084 A.5202.47%i.0928.9072.5188.4812.0909.9091.5222.4778.0912.9088 47 759 2"0.o030 A.i970 A.1613 A.8387 A.601b.J382.i616.8384. 0i41.3959.1611. 8389

METHANE NEO-PENTANE 6INARY SYSTEM METHANE NEO-PENTANL BINARY SYSTEM RUN PR ES S TEMP VAPOR COMPOSITION LIQUID COMPOSITION RUN PRESS TEMP VAPOR COMPOSITION LIQUID COMPOSITION NUMBER PSIA (F) CI NEOC5 CI NEOC5 NUMBER PSIA (F) cl NEOC5 Cl NEOC5 71 511 160.7608 A.2392 A.1532 A.8468 A 75A 1521 160. 7842 A. 2158 A.4815 A.5185 A.7635.23ta5.1535.8465. 7817. 21 83.4822.5178.7580.2420.1529.8471. 7866. 2134.4808.5192 72 763 1bO.7973 A.2027 A.2322 A.7678 A 76B 1709 1bO.7274 A.2726 A.5601 A.4399 A.79 57.2043.2312.7688. 72b1.2739.5595.4405.1990.2010.2 3 31.7669. 728b.2714.5608.4392.2324.7676 77 1748 160.b853 A.3141 A.6028 A.3972 A 73 1005 lbU.8191 A.1809 A.3118 A.6882 A.6835.3165 tb004.3996.8176.1824.3114.6886.6871.3129.6056.3944.82 14.1766.3122.6818.6025.3975.8181.1819 82 310 160.6670 A.3330 A.0852 A.9148 A 74 1213 1 o0.8133 A.1667 A.3909 A.6091 A.6b67.3333.0849.9151. 6IL —4.1876.3914.6086.6672.3328.0855.9145.8142.1858.3904.6096.085~.9149 74A 1281 160.812o A.1i874 A.3975 A.6025 A 83 308 220.3949 A.6051 A.0505 A.9495 A.8126.18ts14.3991.6009.3912.6088.0507.9493.39 59.6041.3986.6014.0503.9497.3949.6051 84 503 220.5631 A.4369 A.1168 A.8832 A.5628.4372.1163.8837,.5616.4384.1168.8832.5650.4350.1172.8828

-141METHANE NEO-PENTANE BINARY SYSTEM RUN PRESS TEM-P VAPOR COMPUSITION LIQUID COMPOSITION NUMBER PSIA F ) C 1 NEOC C NEOC5 85 748 220.6389 A.3611 A.1970 A.8030 A.6374. 3626. 1967.8033.6403.3597.197b. 8024 1966.8034 86 1008 220.6703 A.3297 A.2819 A.7181 A.67 11.3289.2822.1178.6696.3304.2816.7184 87A 1251 220.6535 A.3465 A.3766 A.6234 A. 65 35.3465. 37167.6233.b6535.3465.3765.6235 88A 1434 220.5852 A.4148 A.4712 A.5288 A 58 47.4153.4700.5300. 5857.4143.4724. 52 76 91 500 280.2603 A.7197 A.0683 A.9317 A.2801. 71 99.0688.9312.2b05. 7195. 06 78. 932 2.0684.9316 92 755 280.4066 A.5932 A.1632 A.8368 A.4082.5918.1635.8365.4054.5946.1659.8341.1604.8396 938 1004 280.4159 A.5841 A.2813 A.7187 A.4155.5t45.2817. 7183. 4158.5842.28 1 4. 7 186.4104.5836.2807.7193

-142TABLE XXVII EXPERIMENTAL DATA FOR TERNARY SYSTEMS

METHANE NEUPENTANE NORMAL PENTANE TERNARY SYSTEM M~ETHANE -IEdPr-4FAl,, NJkRiL Pt6NTANE TERNARY SYSTEM RUN PRESS TEMP VAPOR COMPOSITION LIQUID COMPOSITION RUN p K S5 TEIP VAPOR COMPOSITION LIQUID COMPOSITION NUMBER PSIA (F) Cl NEOC5 N-C5 CI NEOC5 N-C5 NUMbEK PS5IMA (F) Cl NEOC5 N-C5 CI INEOC5 N- C5 95 503 160.6451 A.0577 A.0972 A.1407 A.2165 A.6429 A 99 ltb9 1o0.8546 A.0429 A.1025 A.4607 A.1314 A.409A.8438.0582.0980.1409.2165.6426.8566.0422.1012.4595.1318.4086.8464.0511.0964.1405.2164.6431.8512.0438.1050.4619.1310.401.8559.0427.1014 96 751 160.8714 A.0459 A.0827 A.2056 A.1997 A.5945 A.8706.0460.0832.2054.2000.5946 100 2013 160.8052 A.0536 A.1413 A.5500 A.1112 A.3388.8721.0458.0822.2063.1995.5943.803 7.0540.1423.5494.1109.3397.8066.0531.1402.5518.1109.337 - 97 1251 160.8778 A.0404 A.0818 A.3375 A.1641 A.4985 A -.5488.1118.3393.8771.0404.0826.3367.1640.4993.8766.0404.0810.3383.1641.4976 101 2120 160.7753 A.0591.1655-A.6019 A.0976 A.3005.7742.0589.1669.6001.0981.301 98 iSOS 160.8699 A.0410 A.0891 A.3998 A.1482 A.4520 A.7764.0594.1642.6023.0979.2998.8699.0407.0893.3976.1489.4535.6033.0968.2999.8699.0412.0889.4021.1415.4505 102 bOob 160.8187 A.0406 A. 0807 A.2781 A.1760 A.5459.8791.0404.0805.2783.1759.545&.8782.0408.0810.2779.1761.5460

METHANE ISOPENTANE NORMAL PENTANE TERNARY SYSTEM. METHANE ISOPENTANE NORMAL PENTANE TERNARY SYSTEM RUN PRESS TEMP VAPOR COMPOSITION LIQUID COMPOSITION RUN PRESS TEMP VAPOR COMPOSITION LIQUID COMPOSITION NUMBER PSIA (F) C1 ISOC5 N-C5 C1 ISOC5 N-C5 NUMBER PSIA (F) C! ISOC5 N-C5 C1 ISUC5 N-C5.51 ~04 160.8712 A.0395 A.089Z A.1386 A.2227 A.63'87 A 55 1975 160.8488 A ~0405 A.1107 A.5038 A.1290 A.3672 A.8713.0402.0885.1386.2227.6387.8491.0398.1'112.5029.1293.3678.8712.0389.0899.848~.0412.1103.5048.1287 ~3665 52 755 160.89::1,9 A.0305 A.0755 A.2114 A.2058 A.5828 A 55A 199.5 160.8419 A.04ZZ A./159 A.5210 A o1229 A o3561 A.8938.0305.0757.2LO1.2057.5842.8372.0430.1198.5240.1219.35~1 I.8940.0300.0754.2k27.2059.5814.8467.041~.E12U.5180.1239.3580 } —d m 53 1003 lbO.897! A.029b A.0733 A.2738 A.1888 A.5374 A 58 2208 160.7585 A.0632 A.1783 A,5934 A.1039 A.3027 A.8967.0300 ~0733.2723.1888.5389.7566.0639.1795 ~5894.1050.3056.d976.0292.0732.2753.1888.5358,760b.0625,1770.5973,1029.2998 5% 1q93 lb0.8890 A.0307 A.0803 A.4057 A.1536 A.4408 A 59 1765 220.7706 A.0613 A.1681 A.4537 A,1391 A.4072 A.b884.0309.OB07.4039.15]9.442!.768[.Ob19.l?00.4536.1398.4066.8896.030b.0798.4074.[552.4394.773[.Oh07 ~.[662,4539.1384.4078 60 2041 ZZO.?4tO A.0652 A.]878 A.5552 A.~1[? A.BB31 A.?~36.Obb9 o1906.5543.1119.]338. -[505.0645.1849.5561.1115.3324

METHANE ISOPENTANE NORMAL PENTANE TERNARY SYSTEM F.-HANE ISOPENTANE NORMAL PENTANE TERNARY SYSTEM RUN PRESS TEMP _VAPOR COMPOSITION LIQUID COMPOSITIUN RUN PRESS TEMP VAPOR COMPOSITION LIQUID COMPOSITION NUMBER PSIA (F) Cl ISOC5 N-C5 Cl ISOC5 N-CS NUMBER PSIA (F) Cl ISOC5 N-C5 Cl ISOC5 N-C5 61 1519 220.8006 A.0527 A.1466 A.3887 A.1540 A.4573 A 66 539 280.5633 A.1162 A.3205 A.0998 A.2196 A.6806 A.7976.0539.1485.3870.1548.4582.5624.1165.3211.1001.2195.6804.8037.0515.1448.3905.1532.4563.5642.1159.3198.0994.2197.6809 5 2. 1263 220.8100 A.0509 A.1391 A.3191 A.1716 A.5093 A 67 7o0 280.6303 A.0967 A.2730 A.1622 A.2017 A.6361 A.8111.0507.1382.3179.1716.5104.6306.0973.2721.1630.2007.6363.8089.0510.1401.3203.1716.5082.6300.0962.2739.1o14.2028.6358 S3 j995 220.8090 A.0510 A.1394 A.2513 A.1816 A.5611 A 68 1001 280.6623 A.0866 A.2510 A.2311 A.1866 A.5822 A.8073.051s.1410.2417 8.1880.5642.6tj25.0862.2512.2295.1861.5844.8119.0504.1377.2548.18172.5580.b621.0870.2509.2327.1872.5800 )04 753 22u 78 76 A.0571 A.1545 A.1867 A.2047 A.6086 A 69 1253 280.6738 A.0809 A.2452 A.3064 A.1644 A.5291 A.7845.0592.1564.1857.2047.6097.6736.0803.2461.3060.1654.5287.7911.05o3.152o.1878.2047.6075.6740.0ido.2444.3068.1635.5296 I)5 507 220.7'45 A.07 0 A.1854 A.1202 A.2198 A.6599 A 66A 541 280.5682 A.1197 A.3120 A.1030 A.2321 A.6649 A.7400.0725.1875.1201.219j.6605. 5o).1207.3133.1036.2337.6627.7490.0677.1833.1203.2203.6593.5705.1188.3107.1025.2305.6670

METHAiNEl'ISoPENrANLE NORMAL PEINTANE TERKNAKY SYSTEM RUN PR SS it''MP VAPU"R COCMP&JSTILN LIQ3UID CUMPOSITION NUMBER PSIA IF) CLI ISu'C5 N-L5 Ci ISOC5 N-C5 67A 157 280.6292 A.1020 A.2687 A.1o42 A.2143 A.621.5 A.6262.102d.2710.1647.2145.6207.o323.1012.2665.1637.2141.6222 68A 1031 280.6646 A.0910 A.2444 A.2417 A.1.956 A.5628 A H.6625.091.2 4 5.2385.1967.5649.b668.0902.2430.2449.1944.5607 69A 1.2 5 26o0.674! )A.0870 A.2365 A.3042 A.1 7197 A.51 61 A.6679.0864.2436.3070.1.768.5.162.6611.085o 02333.3014.1827.5160 708 1565 280.6156 A.0194 A.2850 A.4533 A.1384 A.4083 A.6156.0994.2850.4533.1384.4033

APPENDIX C CALIBRATIONS A, Calibration of Pressure Gauge The Heise pressure gauge (Model No. H24564) was calibrated using a dead weight tester~ The tester (No. 1315) was supplied by the American Gauge Company. The calibration results are given in Table XXVIIIo TABLE XXVIII CALIBRATION OF PRESSURE GAUGE Heise Pressure Actual Pressure Gauge Reading (psi) UP DOWN 300 299 300 500 500 500 750 748 750 999 998 999 1249 1248 1249 1499 1499 1499 1749 1748 1749 1999 1997 1998 2249 2246 2247 2498 2496 2496 -147

-148Bo Calibration of Thermometer The gas-filled mercury in glass thermometer (Model No. 1704431), supplied by the Taylor Thermometer Company and used to measure equilibrium temperatures, was calibrated by comparison with previously calibrated Princo thermometers. The calibrations are given in Table XXIX. TABLE XXIX CALIBRATION OF THERMOMETER Princo Thermometer Taylor Thermometer No. 253197 Reading 71.2~C 160~F 71.1~C Princo Thermometer Taylor Thermometer No. 503944 Reading 104, 6~C 220~F - 104.4~C 1381 oC 2800F - 137.8~0C Co Calibration of Gas Chromatograph Several synthetic mixtures of methane-normal pentane, methane-isopentane, and methane-neopentane were prepared for calibration of the Perkin-Elmer gas chromatograph. These mixtures were made up in a mixture-blending system which is described in the "Installation and Training Recommendations" of the Mass Spectrometer Model Noo 21-103B, Consolidated Engineering Corporation, Pasadena, California, The mixtures of the three binary systems covered the range of interest for

-149this work. A computer program was written to calculate the number of moles of gas from the P-V-T measurements. The equation used was of the form: P= RTt + va The second virial coefficient, B, was determined from the equation presented by Pitzer and Curl,* Calibrations were performed by introducing a sample of the mixture into the chromatograph in the same manner as described in Chapter VIo Areas under the resulting chromatographic curves were measured, and the area of ratio of methane to the pentane isomer was computed. This resultant area ratio was then plotted as a function of the corresponding known mole ratio, Figure 23 is a plot of the area ratio as a function of mole ratio of methane to n-pentane. Figure 24 presents the inverse ratios as coordinates to better illustrate some of the data, Figures 25 and 26 are exactly analogous to Figures 23 and 24, except that the former are for the methane-isopentane binary mixtures, Figures 27 and 28 are the calibration curves for the methane-neopentane binary system, A best "least squares" fit line was drawn through the points in Figures 23, 25, and 27. These calibration curves showed peak area to be linear with molar concentration of the sample, Mole fractions were then obtained from a normalized equation of the following form for * Pitzer, KoS., and RoF Curl, Jr, "Empirical Equation for the Second Virial Coeffieient", J. Am. Chem, Soc., 79, 2369 (1957)o

-15020 _.5 an I s 1.0 O0 w 0.5 O 0 1.0 2.0 3.0 4.0 4.5 MOLE RATIO(CH4/n-C5) Figure 23. Gas Chromatograph Calibration for Methane-Normal Pentane System on a Normal Pentane Basis.

12 10 0 0 C) 6.0 o 4.0 w~~~~~~~~~~~~~~~~~~~~~ 2.0 0 1.0 2.0 3.0 4.0 MOLE RATIO(n-C5/CH4) Figure 24. Gas Chromatograph Calibration for Methane-Normal Pentane System on a Methane Basis.

-1521.7 1.5 0 - cn U) 0 1.0 wcr 0.5 O C I I. I I I 0 1.0 2.0 3.0 4.0 MOLE RATIO (CH4/ ISO C5) Figure 25. Gas Chromatograph Calibration for Methane-Isopentane System on an Isopentane Basis,

-1537.0 6.0I 0 5.0 0 cn 0 4.0,w 3.0 2.0 1.0 o 0, I I, I i I I I I I I I o 1.0 2.0 3.0 4.0 MOLE RATIO(ISO C5/CH4) Figure 26. Gas Chromatograph Calibration for Methane-Isopentane System on a Methane Basis.

-1541.7 1.5 0 O 1.0 2.0 3.0 4.0 Ir 0.5 MO0 1.0 2.0 3.0 4.0 MOLE RATIO(CH4/NEO C.) Figure 27, Gas Chromatograph Calibration for Methane-Neopentane System on a Neopentane Basis,

12.0 I0.0 o 8.0 z 6. - Q 4.0 w cr 2.0 0 1.0 2.0 3.0 4.0 4.5 MOLE RATIO(NEO C5/CH4) Figure 28. Gas Chromatograph Calibration for Methane-Neopentane System on a Methane Basis,

-156binary mixtures: An equation of the following form was used for composition determination of ternary mixtures: / 4, + fL 2 +c/' 3 where Al is the peak area for methane, and A2 and A3 are the peak areas for the pentane isomers, The symbols c and c' are the relative response factors established from the slope of the calibration curves. These response factors correct for the difference of thermal conductivity of the components. Table XXX compares the methane composition as determined from the analytical technique with the known mole fraction for the three prepared mixtures of methane-n-pentane. Several analyses were made for each mixture. Table XXXI compares the sample composition determined by chromotography with the known composition for four prepared methaneisopentane mixtures. Table XXXII compares chromatographic composition determinations with known methane-neopentane mixtures. Included in Tables XXX, XXXI, and XXXII are analyses of three mixtures determined with the aid of a mass spectrometer, Only one mixture of each of the three binary systems was subjected to mass spectrometer analysis.

-157TABLE XXX COMPARISON OF ANALYSES FOR METHANE-N-PENTANE MIXTURES Actual Composition G.C, Analyses MS. Analyses mole % mole % mole % Blend CH4 n-C5 CH4 n-C5 CH4 n-C5 19o 43 80.57 19o41 80.59 19.88 80,12 19.49 80.51 19.50 8o, 50 19,53 80,47 19o25 80,75 Avg,,_19,48 80.52 19o54 80,46 2 80,1 19,9 80,0 20.0 80,l 19 9 80,1 19,9 AvE 80,1 19,9 3 56,1 43~9 56,0 44,0 56,1 43.9 56,2 43,8 55.8 44,2 Avg, 56,0 44,0

-158TABLE XXXI COMPARISON OF ANALYSES FOR METHANE-ISOPENTANE MIXTURES Actual Composition G.C. Analyses M.S. Analyses mole % mole % mole % Blend CH4 i-C5 CH4 i-C5 CH4 i-C5 1 30.4 69.6 30.5 69.5 30.5 69.5 30.4 69.6 Avg. 30.5 69.5 2 34.4 65.6 34.4 65.6 35.14 64.86 34.7 65.3 35.01 64.99 34.4 65.6 33.84 66.16 Avg. 34.5 65.5 34.67 65.33 3 52.1 47.9 51.9 48.1 52.0 48.0 52.6 47.4 52.1 47.9 Avg. 52.2 47.8 4 79.7 20.3 79.8 20.2 79.6 20.4 79.8 20.2 79.7 20.3 Avg. 79.7 70.3 5 23.4 76.6 23.8 76.2 23.6 76.4 23.7 76.3 23.3 76.7 23.7 76.3 Avg. 23.6 76.4

-159TABLE XXXII COMPARISON OF ANALYSES FOR METHANE-NEOPENTANE MIXTURES Actual Composition G.C. Analyses M.S. Analyses mole % mole % mole % Blend CH4 neo C5 CH4 neo C5 CH4 neo C5 1 19.14 80.86 19.49 80.51 19.47 80.53 19.57 80.43 Avg. 19.51 80.49 2 26.8 73.2 26.7 73.3 26.51 73.49 26.7 73.3 26.27 73.73 26.3 73.7 26.47 73.53 26.5 73.5 Avg. 26.5 73.5 26.41 73.59 3 79.2 20.8 79.2 20.8 79.1 20.9 79.5 20.5 79.3 20.7 Au. 79.4 20.6 4 66.7 33.3 66.8 3355.2 66.6 33.4 66.7 33.3 66.8 33.2 66.8 33.2 Avg. 66.8 33.2

APPENDIX D GRAPHICAL COMPARISONS OF CALCULATED K-VALUES WITH OBSERVED K-VALUES Figures 29 through 38 present comparisons of the calculated equilibrium vaporization ratios with the observed equilibrium vaporization ratios as a function of pressure. The experimental or observed K-values, are represented by a solid curve and the calculated values, as presented in Table XXV, are represented by a dashed curve. Visual inspection of Figures 29 through 38 indicate that the analytical expressions does not adequately represent the phase behavior of methane in enopentane at high temperatures.(see Figure 34.) -160

IO 8.0 - EXPERIMENTAL VALUE - -__ CALCULATED VALUE 6.0 -1 5.0 4.0 3.0 2.0 c 1.0 M 0.8 0.6 j o 0.5- ii 0.4 - 0.3 0.2 - ic5 - 0.2 0.1.. I I I I I,I II 100 200 300 500 700 1000 2000 3000 5000 PRESSURE (PSIA) Figure 29. Comparison of Calculated K with Observed K for Methane-Isopentane Binary at 160 F.

-1621 0 8.0 - EXPERIMENTAL VALUE --.. CALCULATED VALUE 6.0 C1 5.0 4.0 3.0 2.0 Y O C 1.0 D 0.8 0.60.4 / 0.3 0.2 0.1I I I I I. 100 200 300 500 700 1000 2000 3000 5000 PRESSURE (PSIA) Figure 30. Comparison of Calculated K with Observed K for Methane-Isopentane Binary at 220 ~F.

-163I 0 8.0 - EXPERIMENTAL VALUE... _ CALCULATED VALUE 6.0 C 5.0 4.0 3.0 2.0 CC 1.0 2 0.8 Z 0.6 iC~ " 7' O 0.5 " 0.4 0.3 0.2 0.1 L, I,,! I I 1 l 100 200 300 500 700 1000 2000 3000 5000 PRESSURE (PSIA) Figure 31. Comparison of Calculated K with Observed K for Methane-Isopentane Binary at 280 ~F.

0 &O - C1 EXPERIMENTAL VALUE - \__CALCULATED VALUE 6.0 5.0 4.0 3.0 2.0 C 1.0 2 0.8 - 0.6 00.5 0.4~~ neo ic5. ii 0.3 0.2 0.1 I I I I 100 200 300 500 700 1000 2000 3000 5000 PRESSURE (PSIA) Figure 32. Comparison of Calculated K with Observed K for Methane-Neopentane Binary at 160~F.

-165I0 8.0 - C1 EXPERIMENTAL VALUE -... CALCULATED VALUE 6.0 5.04.0 3.0 2.0 0 C 1.0 2 0.8 mC P LIC)0.6 neo C5 ==X 0.5' -/ 0.4 - 0.3 0. 2 0.1 I I i I 100 200 300 500 700 1000 2000 3000 5000 PRESSURE (PSIA) Figure 33. Comparison of Calculated K with Observed K for Methane-Neopentane Binary at 2200F.

-166I 0 8.0 - EXPERIMENTAL VALUE - -.CALCULATED VALUE 6.0 5.0 4.0 3.0 \ " 2.0 1.0 2 0.8 neo C5 0.6 o 0.5 0.40.3 0.2 0.1 I I I I 100 200 300 500 700 1000 2000 3000 5000 PRESSURE (PSIA) Figure 34. Comparison of Calculated K with Observed K for Methane-Neopentane Binary at 280 ~F.

-167I0 8.0 - EXPERIMENTAL VALUE.... _ CALCULATED VALUE 6.0 - Cl 5.0 "' 4.0 3.0 2.0 1.0 0.8 1 0.6 o 0.5O.40.3 neo c5.,C _ / 0.2 /n-CO =... 2L l 0.1 I I 100 200 300 500 700 1000 2000 3000 5000 PRESSURE (PStA) Figure 35. Comparison of Calculated K with Observed. K for Methane-NeopentaneNormal Pentane Ternary at 160~0F,

-16810. 8.0 - EXPERIMENTAL VALUE ___ CALCULATED VALUE 6.0 C1l 5.0 4.0 3.0 - 2.00: i.0 0.8 J 0.6 ~ O.5w 0.4 0.3 0.2 iC < l.0..1 I -l.1, I -nC,. I 100 200 300 500 700 1000 2000 3000 5000 PRESSURE (PSIA) Figure 36. Comparison of Calculated K with Observed K for Methane-IsopentaneNormal Pentane Ternary at 160~F.

-169I0 8.0 - EXPERIMENTAL VALUE - C-.... CALCULATED VALUE 6.0 - 5.04.0 - 3.0 - 2.0 1.00.8 co, -JO.6 0 / iCi 0.3 - nC 0.2 0.1 I I I III 100 200 300 500 700 1000 2000 3000 5000 PRESSURE (PSIA) Figure 37. Comparison of Calculated K with Observed K for Me-thane-IsopentaneNormal Pentane Ternary at 2200F.

UNIVERSITY OF MICHIGAN II III IIIi HIII111lllil3 9015 03695 1633 -170I0 8.0 - EXPERIMENTAL VALUE..... CALCULATED VALUE 6.0 C1 5.04.0 - 3.0 2.0 C 1.0 o 0.8 O 0.6 iC5 0.5 nC5 0.4 0.3 0. 2 0.1, 100 200 300 500 700 1000 2000 3000 5000 PRESSURE (PSIA) Figure 38. Comparison of Calculated K with Observed K for Methane-IsopentaneNormal Pentane Ternary at 280 F.