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Satyro M.A.,Virtual Materials Group Inc. | Shaw J.M.,University of Alberta | Yarranton H.W.,University of Calgary
Fluid Phase Equilibria | Year: 2013

A simple model is proposed for the calculation of oil and water mutual solubilities as a function of temperature using the hydrocarbon specific gravity and Watson-K factor as correlating parameters. The model parameters were determined using a single consistent set of high quality solubility data collected at the water or hydrocarbon saturation pressures at temperatures from 273. K and 573. K but primarily from 298. K to 398. K. The hydrocarbon in water solubilities ranged over 11 orders of magnitude while the water in hydrocarbons solubilities ranged over 4 orders of magnitude. Hydrocarbon types ranged from low molecular weight paraffins such as n-pentane to heavy aromatics such as anthracene. The average absolute error in the estimated solubilities of water in hydrocarbons over 621 data points was 34% and approached the suggested underlying uncertainty in the reference data (30%). The average absolute error in the estimated solubilities of hydrocarbons in water over 964 data points was 88% and exceeded the suggested underlying uncertainty in the reference data (30%).The model can be extended to higher temperatures using a procedure based on activity coefficients or equations of state and is easily adapted to work with ill-defined hydrocarbon mixtures. Prediction and correlation of solubilities of water in kerosene, heavy oils, and bitumen were made at temperatures, pressures, and compositions far removed from the original data used for model construction. The correlation failed near the critical temperature of water because under these conditions, the properties hydrocarbon-rich phases in the mixtures on which the correlation is based (Type II and Type IIIa phase behaviour according to the van Konynenburg and Scott naming scheme) and those of water. +. heavy hydrocarbons (Type IIIb phase behaviour) diverge. A separate NRTL based correlation is presented for the mutual solubility of water. +. heavy oils at temperatures close to the critical temperature of water. © 2013 Elsevier B.V. Source


Castellanos-Diaz O.,University of Calgary | Schoeggl F.F.,University of Calgary | Yarranton H.W.,University of Calgary | Satyro M.A.,Virtual Materials Group Inc.
Industrial and Engineering Chemistry Research | Year: 2013

The prediction of heavy oil phase behavior, particularly with solvents, is sensitive to the characterization of the middle and heavy boiling point components of the oil. These components are typically characterized based on an extrapolation of distillation data. One method to test the extrapolated characterization is to model the vapor pressures of these fractions or residues containing these fractions. Unfortunately, the vapor pressures are too low to be reliably measured with conventional techniques. A new high vacuum static apparatus was designed and constructed for the measurement of vapor pressure of heavy oil and bitumen samples. The apparatus is capable of measuring pressures from 100 down to 0.1 Pa and temperatures in the range of 293.15-473.15 K. New procedures were developed to degas samples and obtain accurate vapor pressures at vacuum conditions. The apparatus was tested on n-hexadecane and naphthalene at temperatures between 303.15 and 363.15 K. The measured vapor pressures were, on average, all within 13% of the literature data. The vapor pressures of a Western Canadian bitumen sample (WC-BIT-B1) and three of its fractions were measured using the apparatus. The WC-BIT-B1 bitumen was modeled using the Advanced Peng-Robinson equation of state using a Gaussian extrapolation of its distillation curve for the maltene fraction and a Gamma molecular distribution for its asphaltene fraction. The measured vapor pressures were all predicted to within 3.5%. © 2013 American Chemical Society. Source


Hajipour S.,University of Calgary | Satyro M.A.,Virtual Materials Group Inc. | Foley M.W.,University of Calgary
Fluid Phase Equilibria | Year: 2014

A simple procedure is proposed to evaluate the uncertainty on binary interaction parameters propagated from the uncertainties present in the physical properties and equation parameters applied for their calculation using a simple cubic equation of state. A useful database containing 87 binary mixtures present in natural gas processing was constructed through critical collection of available experimental vapour-liquid equilibrium (VLE) data and their associated uncertainties.A thermodynamic consistency test was performed on each isothermal dataset to investigate the quality of the associated VLE data. Upon acceptance of the VLE data based on its quality and consistency, binary interaction parameters and associated uncertainties were determined using a combination of nonlinear regression and Monte Carlo simulation, taking into consideration the uncertainties of pure components, equation of state parameters, and VLE data. The Monte Carlo simulation was also used for the error propagation to estimate the uncertainty on the calculated VLE. Sample calculations were presented illustrating the effect of uncertainties on the PXY and TXY diagrams of ethane/propane and methane/hydrogen sulfide binary mixtures. The required minimum number of stages for a simplified de-ethanizer was calculated taking into account uncertainties of basic input parameters. In addition the effect of uncertainties on the position of calculated cricondenbar and cricondentherm was investigated. © 2013 Elsevier B.V. Source


Satyro M.A.,Virtual Materials Group Inc. | Satyro M.A.,University of Calgary | Schoeggl F.,University of Calgary | Yarranton H.W.,University of Calgary
Fluid Phase Equilibria | Year: 2011

In natural gas dehydration units, rich TEG solutions are decompressed before the TEG regeneration stage and the direction of the temperature change during the decompression has been debated. The temperature change from an isenthalpic expansion from (7000. kPa to 440. kPa was measured for the following aqueous mixtures: pure water, 99% pure triethylene glycol (TEG), aqueous TEG (99. wt% TEG. +. 1% water), aqueous TEG saturated with methane, aqueous TEG saturated with n-pentane, and aqueous TEG saturated with n-heptane. In all cases, the temperature increased upon expansion with the magnitude of the temperature change ranging from 1.4. K for pure water to 2.4. K for TEG. A simple equation of state model predicted the correct direction for the temperature change and the predicted values were within ±1. K of the experimental data. © 2011 Elsevier B.V. Source


Mortazavi-Manesh S.,University of Calgary | Satyro M.,Virtual Materials Group Inc. | Marriott R.A.,University of Calgary
Canadian Journal of Chemical Engineering | Year: 2013

Solubility information for CO2 in different ionic liquids, ILs, in part can potentially be used to select a specific IL for the separation of CO2 from hydrocarbon fluids. Unfortunately, not all CO2-IL systems have been experimentally described at similar temperatures and pressures; therefore, a direct comparison of performance by process simulation is not always possible. In the extreme cases, the design of a CO2 separation process may require predicting the CO2-IL equilibria for which there are no available solubility data. To address the need for this information, a semi-empirical correlation was developed to estimate the dissolution of CO2 in CO2-IL solvent systems. The theoretical COSMO-RS calculation method was used to calculate the chemical potential of CO2 in a wide variety of ILs and the Soave-Redlich-Kwong equation was used to calculate the fugacity coefficient of the CO2 vapour phase. The model was correlated with available literature data, yielding an average error of AAR=23% and small bias. © 2012 Canadian Society for Chemical Engineering. Source

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