Equilibrium system nonideal vapor/liquid

Ideal Vapor/Liquid Equilibrium Systems, Nonideal Vapor/Liquid Equilibrium Systems, Vapor/Liquid Equilibrium Relationships,... 

An extensive tabulation of azeotropes has been compilnd by Horsley.1 An older compilation is that of Lecat.2 Certain nonideal vapor-liquid equilibrium models ate useful for ptediciiag azeotropic behavior of binery systems in particular, the model of Rcoon and Prausnitz3 is usefol in this regard because it can handle the two liquid phases associated with heterogenenus azeotropes. The Horsley book also contains guidelines for the prediction of azeotropes. 

To account for nonideal vapor-liquid equilibrium and possible VLLE for this quaternary system, the NRTL model is used to calculate the activity coefficients. Model parameters are taken from Chapter 7. Vapor-phase nonideality caused by the dimerization of acetic acid is also taken into consideration using the Hayden-O Connell second virial coefficient model. Aspen Plus built-in parameters values are used. 

The compositions of the vapor and liquid phases in equilibrium for partially miscible systems are calculated in the same way as for miscible systems. In the regions where a single liquid is in equilibrium with its vapor, the general nature of Fig. 13.17 is not different in any essential way from that of Fig. I2.9< Since limited miscibility implies highly nonideal behavior, any general assumption of liquid-phase ideality is excluded. Even a combination of Henry s law, valid for a species at infinite dilution, and Raoult s law, valid for a species as it approaches purity, is not very useful, because each approximates real behavior only for a very small composition range. Thus GE is large, and its composition dependence is often not adequately represented by simple equations. However, the UNIFAC method (App. D) is suitable for estimation of activity coefficients. 

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. 

This study was undertaken to obtain the necessary vapor-liquid equilibrium data and to determine the distillation requirements for recovering solvent for reuse from the solvent-water mixture obtained from adsorber regeneration. Previous binary vapor-liquid equilibrium data (2, 3) indicated two binary azeotropes (water-THF and water-MEK) and a two phase region (water-MEK). The ternary system was thus expected to be highly nonideal. 

Gas solubility has been treated extensively (7). Alethods for the prediction of phase equilibria and actual solubility data have been given (8,9) and correlations of the equilibrium K values of hydrocarbons have been developed and compiled (10). Several good sources for experimental information on gas— and vapor—liquid equilibrium data of nonideal systems are also available (6,11,12). 

We conclude this discussion with one final reminder. The vapor-liquid equilibrium calculations we have shown in Section 6.4c are based on the ideal-solution assumption and the corresponding use of Raoult s law. Many commercially important systems involve nonideal solutions, or systems of immiscible or partially miscible liquids, for which Raoult s law is inapplicable and the Txy diagram looks nothing like the one shown for benzene and toluene. 

It can be seen from equation (34.18) that for an ideal liquid system, the vapor will always be relatively richer than the liquid in the more volatile constituent, i.e., the one with the higher vapor pressure. For example, if, as in Fig. 21, the component 1 is the more volatile, p will be greater than p, and hence Ni/i 2 will exceed Ni/ns, and the vapor will contain relatively more of the component 1 than does the liquid with which it is in equilibrium. This fact is fundamental to the separation of liquids by fractional distillation. The limitations arising in connection with nonideal systems will be considered in 35b, 35c. 

Various processes are used for separating components that are difficult or impossible to be separated by conventional distillation. Whether the difficulty of separation arises from the components close boiling points or their tendency to form azeotropes, the separation processes must take into account the complex vapor-liquid equilibrium relationships of the system. The system to be considered involves both the components to be separated and the separating agent that, in one way or another, enhances the desired separation. The vapor-liquid equilibria of such mixtures is highly nonideal, and it is precisely this nonideality that is capitalized on to bring about the separation. 

As the first illustration of the use of these equations, consider vapor-liquid equilibrium in the hexane-triethylamine system at 60°C. These species form an essentially ideal mixture. The vapor pressure of hexane af this temperature is 0.7583 bar and that of triethylamine is 0.3843 bar these are so low that the fugacity coefficients at saturation and for the vapor phase can be neglected. Consequently, Eqs. 10.1-3 and 10.1-4 should be applicable to this system. The three solid lines in Fig. 10.1-1 represent the two species partial pressures and the total pressure, which were calculated using these equations and all are linear functions of the of liquid-phase mole fraction the points are the experimental results. The close agreement between the computations and the laboratory data indicates that the hexane-triethylamine mixture is ideal at these conditions. Note that this linear dependence of the partiaLand total pressures on mole fractions predicted by Eqs. 10.1-2 and 10.1-3 is trae only for ideal mixtures it is not true for nonideal mixtures, as we shall see in Sec. 10.2. 

Few liquid mixtures are ideal, so vapor-liquid equilibrium calculations can be more complicated than is the case for the hexane-triethylamine system, and the system phase diagrams can be more structured than Fig. 10.1-6. These complications arise from the (nonlinear) composition dependence of the species activity coefficients. For example, as a result of the composition dependence of y, the equilibrium pressure in a fixed-temperature experiment will no longer be a linear function of mole fraction. Thus nonideal solutions exhibit deviations from Raoult s law. We will discuss this in detail in the following sections of this chapter. However, first, to illustrate the concepts and some of the types of calculations that arise in vapor-liquid equilibrium in the simplest way, we will assume ideal vapor and liquid solutions (Raoult s law) here, and then in Sec. 10.2 consider the calculations for the more difficult case of nonideal solutions.. ... 

Construction of Vapor-Liquid Equilibrium Diagrams for a Nonideal System... 

As another example of low-pressure vapor-liquid equilibrium, we consider the n-pentane-propionaldehyde mixture at 40.0 C. Eng and Sandler took data on this system using the dynamic still of Fig. 10.2-5. The x-y-P-T data in Table 10.2-1 and Fig. 10.2-8fl and b were obtained by them. (Such data can be tested for thermodynamic consistency see Problem 10.2-12.) As is evident, this system is nonideal and has an azeotrope at about 0.656 mole fraction pentane and 1.3640 bar. We will use these data to test the UNIFAC prediction method. 

The dashed lines in the figure are the predictions, at all temperatures for the acetone-water system that result from setting the binary parameter k 2 equal to zero. Note that very nonideal behavior is predicted, which shows that setting ]c 2 = 0 is not equivalent to assuming ideal solution behavior. In fact, such extreme nonideal behavior is. predicted that vapor-liquid equilibrium calculations made with the program VLMU do not even converge for acetone mole fractions less than about 0.15. 

Figure 11.3-3 shows the vapor-liquid and liquid-liquid equilibrium behavior computed for the system of methanol and n-hexane at various temperatures. Note that two liquid phases coexist in equilibrium to temperatures of about 43°C. Since liquids are relatively incompressible, the species liquid-phase fugacities are almost independent of pressure (see Illustrations 7.4-8 and 7.4-9), so that the liquid-liquid behavior is essentially independent of pressure, unless the pressure is very high, or low enough for the mixture to vaporize (this possibility will be considered shortly). The vapor-liquid equilibrium curves for this system at various pressures are also shown in the figure. Note that since the fugacity of a species in a vapor-phase mixture is directly proportional to pressure, the VLE curves are a function of pressure, even though the LLE curves are not. Also, since the methanol-hexane mixture is quite nonideal, and the pure component vapor pressures are similar in value, this system exhibits azeotropic behavior. 

In the example distillation system considered in Chapters 3 and 4, we studied the binary propane/isobutane separation in a single distillation column. This is a fairly ideal system from the standpoint of vapor-liquid equilibrium (VLE), and it has only two components, a single feed and two product streams. In this chapter, we will show that the steady-state simulation methods can be extended to multicomponent nonideal systems and to more complex column configurations. 

Up to this point we have looked at systems with fairly ideal vapor-liquid equilibrium behavior. The last separation system examined is a highly nonideal ternary system of methyl acetate (MeAc), methanol (MeOH), and water. Methyl acetate and methanol form a homogeneous minimum-boiling azeotrope at 1.1 atm with a composition of... 

If the top temperature is too cold and the bottom tenperature is too hot to allow sandwich conponents to exit at the rate they enter the column, they become trapped in the center of the column and accumulate there fKister. 20041. This accumulation can be quite large for trace conponents in the feed and can cause column flooding and development of a second liquid phase. The problem can be identified from the simulation if the engineer knows all the trace conponents that occur in the feed, accurate vapor-liquid equilibrium (VLE) correlations are available, and the simulator allows two liquid phases and one vapor phase. Unfortunately, the VLE may be very nonideal and trace conponents may not accumulate where we think they will. For example, when ethanol and water are distilled, there often are traces of heavier alcohols present. Alcohols with four or more carbons (butanol and heavier) are only partially miscible in water. They are easily stripped from a water phase (relative volatility 1), but when there is litde water present they are less volatile than ethanol. Thus, they collect somewhere in the middle of the column where they may form a second liquid phase in which the heavy alcohols have low volatility. The usual solution to this problem is to install a side withdrawal line, separate the intermediate component from the other components, and return the other components to the column. These heterogeneous systems are discussed in more detail in Chapter 8. 

The vapor-liquid equilibrium for the very nonideal systems studied in this chapter may not be fit well with any of the correlations in Aspen Plus if the parameters embedded in Aspen Plus are used. An alternative is to use Aspen Plus to fit the parameter values to give the best fit to VLE data. This procedure is explained in Appendix B at the end of the book. 

---