Vapor-liquid equilibria experimental determination

The vapor-liquid equilibrium of the binary mixture is well fitted by Van Laar s equations (228). It was determined from 100 to 760 mm Hg. and the experimental data was correlated by the Antoine equation (289, 290), with P in mm Hg and t in °C ... 

It is essential to calculate, predict or experimentally determine vapor-liquid equilibrium data in order to adequately perform distillation calculations. These data need to relate composition, temperature, and system pressure. 

These models are semiempirical and are based on the concept that intermolecular forces will cause nonrandom arrangement of molecules in the mixture. The models account for the arrangement of molecules of different sizes and the preferred orientation of molecules. In each case, the models are fitted to experimental binary vapor-liquid equilibrium data. This gives binary interaction parameters that can be used to predict multicomponent vapor-liquid equilibrium. In the case of the UNIQUAC equation, if experimentally determined vapor-liquid equilibrium data are not available, the Universal Quasi-chemical Functional Group Activity Coefficients (UNIFAC) method can be used to estimate UNIQUAC parameters from the molecular structures of the components in the mixture3. 

Since all experimental data for vapor-liquid equilibrium have some experimental uncertainty, it follows that the parameters obtained from data reduction are not unique3. There are many sets of parameters that can represent the experimental data equally well, within experimental uncertainty. The experimental data used in data reduction are not sufficient to fix a unique set of best parameters. Realistic data reduction can determine only a region of parameters2. 

A procedure is presented for correlating the effect of non-volatile salts on the vapor-liquid equilibrium properties of binary solvents. The procedure is based on estimating the influence of salt concentration on the infinite dilution activity coefficients of both components in a pseudo-binary solution. The procedure is tested on experimental data for five different salts in methanol-water solutions. With this technique and Wilson parameters determined from the infinite dilution activity coefficients, precise estimates of bubble point temperatures and vapor phase compositions may be obtained over a range of salt and solvent compositions. 

Experimental vapor-liquid-equilibrium data for benzene(l)/n-heptane(2) system at 80°C (176°F) are given in Table 1.8. Calculate the vapor compositions in equilibrium with the corresponding liquid compositions, using the Scatchard-Hildebrand regular-solution model for the liquid-phase activity coefficient, and compare the calculated results with the experimentally determined composition. Ignore the nonideality in the vapor phase. Also calculate the solubility parameters for benzene and n-heptane using heat-of-vaporization data. 

The experimental procedures are quite similar to and often confused with pervaporation. The main difference between VMD and pervaporation is the nature of the membrane used, which plays an important role in the separations. While VMD uses a porous hydrophobic membrane and the degree of separation is determined by vapor-liquid equilibrium conditions at the membrane-solution interface, pervaporation uses a dense membrane and the separation is based on the relative perm-selectivity and the diffusivity of each component in the membrane material. 

Pequemh, A., Carlos Asensi, J., Gomis V. Experimental determination of quaternary and ternary isobaric vapor-liquid-liquid equilibrium and vapor-liquid equilibrium for the systems water-ethanol-hexane-toluene and water-hexane-toluene at 101.3 kPa. J. Chem. Eng. Data, 2010, 56, 3991. 

In using simulation software, it is important to keep in mind that the quality of the results is highly dependent upon the quahty of the liquid-liquid equilibrium (LLE) model programmed into the simulation. In most cases, an experimentally vmidated model will be needed because UNIFAC and other estimation methods are not sufficiently accurate. It also is important to recognize, as mentioned in earlier discussions, that binary interaction parameters determined by regression of vapor-liquid equilibrium (VLE) data cannot be rehed upon to accurately model the LLE behavior for the same system. On the other hand, a set of binary interaction parameters that model LLE behavior properly often will provide a reasonable VLE fit for the same system—because pure-component vapor pressures often dominate the calculation of VLE. 

EPAR ATION and purification processes account for a large portion of the design, equipment, and operating costs of a chemical plant. Further, whether or not a mixture forms an azeotrope or two liquid phases may determine the process flowsheet for the separations section of a chemical plant. Most separation processes are contact operations such as distillation, gas absorption, gas stripping, and the like, the design of which requires the use of accurate vapor-liquid equilibrium data and correlating models or, in the absence of experimental data, of accurate predictive methods. Phase behavior, especially vapor-Uquid equilibria, is important in the design, development, and operation of chemical processes. 

The fugacities f or activity coefficients of a liquid solution are measured in vapor-liquid equilibrium experiments. In commonly employed methods, the liquid solution is brought in contact and kept in contact with a vapor mixture of the same components until equilibrium is attained between the phases. A sample of the vapor is then withdrawn and analyzed to determine its mole fractions y, i = 1,2,. Similarly for the liquid sample, the mole fractions X = 1,2,... are determined. Together with the measured p, an experimental point of vapor-liquid equilibrium is given by... 

Consequently, by a regression analysis of very large quantities of activity coefficient (or, as we will see in Sec. 10.2, actually vapor-liquid equilibrium) data, the binary parameters Onm Omn for many group-group interactions can be determined. These parameters can then be used to predict the activity coefficients in mixtures (binary or multicomponent) for which no experimental data are available. 

The illustrations of this section were meant to demonstrate how one can determine activity coefficients from measurements of temperature, pressure, and the mole fractions in both phases of a vapor-liquid equilibrium system. An alternative procedure is at constant temperature, to measure the total equilibrium pressure above liquid mixtures of known (or measured) composition. This replaces time-consuming measurements of vapor-phase compositions with a more detailed analysis of the experimental data and more complicated calculations. ... 

At the lowest pressure in the figure, P = 0.133 bar, the vapor-liquid equilibrium curve intersects the liquid-liquid equilibrium curve. Consequently, at this pressure, depending on the temperature and composition, we may have only a liquid, two liquids, two liquids and a vapor, a vapor and a liquid, or only a vapor in equilibrium. The equilibrium state that does exist can be found by first determining whether the composition of the liquid is such that one or two liquid phases exist at the temperature chosen. Next, the bubble point temperature of the one or either of the two liquids present is determined (for example, from experimental data or from known vapor pressures and an activity coefficient model calculation). If the liquid-phase bubble point temperature is higher than the temperature of interest, then only a liquid or two liquids are present. If the bubble point temperature is lower, then depending on the composition, either a vapor, or. a vapor and a liquid are present. However, if the temperature of interest is equal to the bubble point temperature and the composition is in the range in which two liquids are present, then a vapor and two coexisting liquids will be in equilibrium. 

Using the proposed procedure in conjunction with literature values for the density (11) and vapor pressure (12) of solid carbon dioxide, the solid-formation conditions have been determined for a number of mixtures containing carbon dioxide as the solid-forming component. The binary interaction parameters used in Equation 14 were the same as those used previously for two-phase vapor-liquid equilibrium systems (6). The value for methane-carbon dioxide was 0.110 and that for ethane-carbon dioxide was 0.130. Excellent agreement has been obtained between the calculated results and the experimental data found in the literature. As shown in Figure 2, the predicted SLV locus for the methane-carbon... 

This value is 31.3% lower than the experimental value. This result indicates that the assumption of temperature independence of (A,y-A,() and ViJvjL is not valid. When temperature dependence is taken into account, predictions are improved. For best accuracy, however, Nagata and Yamada" showed that Wilson parameters should be determined by simultaneous fit of vapor-liquid equilibrium and heat of mixing data. 

The set of Equations 25.42-25.48 can be solved provided the following information is available vapor-liquid equilibrium data, for example, the ternary equilibrium data for a typical esterification reaction mass and enthalpy balances around the feed point, reflux inlet, and reboiler to account for the flow rates, compositions, and thermal conditions of the external streams mass transfer coefficients in the absence of reaction (either by experimental determination or estimation from available correlations) liquid holdup (usually from available correlations) and an expression for the reaction rate. Then the equations can be solved by any convenient method, preferably the Runge-Kutta routine, to get the mole fraction of each component as a function of height. 

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