Saitzew


The calculation of single-stage equilibrium separations in multicomponent systems is implemented by a series of FORTRAN IV subroutines described in Chapter 7. These treat bubble and dewpoint calculations, isothermal and adiabatic equilibrium flash vaporizations, and liquid-liquid equilibrium "flash" separations. The treatment of multistage separation operations, which involves many additional considerations, is not considered in this monograph.  [c.6]

The critical temperature of methane is 191°K. At 25°C, therefore, the reduced temperature is 1.56. If the dividing line is taken at T/T = 1.8, methane should be considered condensable at temperatures below (about) 70°C and noncondensable at higher temperatures. However, in process design calculations, it is often inconvenient to switch from one method of normalization to the other. In this monograph, since we consider only equilibria at low or moderate pressures in the region 200-600°K, we elect to consider methane as a noncondensable component.  [c.59]

The most frequent application of phase-equilibrium calculations in chemical process design and analysis is probably in treatment of equilibrium separations. In these operations, often called flash processes, a feed stream (or several feed streams) enters a separation stage where it is split into two streams of different composition that are in equilibrium with each other.  [c.110]

The product streams can be a vapor and a liquid or two immiscible liquids. The process may consist of a single equilibrium stage, as with a flash drum or single mixer-settler, or of a cascade of equilibrium stages, as with staged distillation or extraction. Multistage separations are usually arranged with countercurrent flow of the phases between the stages. The calculation of all such equilibrium separations is based on enthalpy and component material balances over the separation stage, in combination with the requirements that, for each component, the fugacities be equal in the two streams exiting the stage.  [c.110]

In this chapter we present efficient calculation procedures for single-stage equilibrium separations subroutines implementing these procedures are given in Appendices F and G. While we recognize the great importance of multistage separations, it must be realized that the efficient computation of such processes requires very careful resolution of the large number of simultaneous equilibrium stages involved in a countercurrent cascade. The dominant consideration in such multistage computation procedures is usually the technique used to achieve this simultaneous solution rather than the equilibrium treatment of the stages themselves. (Goldstein and Stanfield, 1970 Holland,  [c.110]

The single-stage separations for which we present computational procedures are the incipient separations (one product phase present in very small amount) represented by bubble and dew-point calculations, vapor-liquid equilibrium separations at fixed pressure under isothermal or adiabatic conditions, and liquid-liquid equilibrium separations at fixed pressure and temperature. These calculations are implemented by FORTRAN IV subroutines designed to minimize the number of vapor and liquid-phase fugacity evaluations necessary to achieve satisfactory solutions. This criterion for efficiency of the algorithms is based on the recognition that, with relatively rigorous thermodynamic methods such as those used here, most of the computation effort in any separation calculation is devoted to evaluation of thermodynamic equilibrium functions. It is important to avoid unnecessary calculations of fugacities or fugacity (activity) coefficients in computer programs used in chemical engineering practice.  [c.111]

Smith, B. D., "Design of Equilibrium Stage Processes," McGraw-Hill, New York (1963).  [c.129]

This text will attempt to develop an understanding of the concepts required at each stage during the creation of a chemical process design.  [c.3]

Building an irreducible structure. The first approach follows the onion logic, starting the design by choosing a reactor and then moving outward by adding a separation and recycle system, and so on. At each layer we must make decisions based on the information available at that stage. The ability to look ahead to the completed design might lead to different decisions. Unfortunately, this is not possible, and instead, decisions must be based on an incomplete picture.  [c.8]

This text concentrates on developing an understanding of the concepts required at each stage of the chemical process design. Such understanding is a vital part of process design, whichever approach is followed.  [c.13]

The lack of suitable catalysts is the most common reason preventing the exploitation of novel reaction paths. At the first stage of design, it is impossible to look ahead and see all the consequences of choosing one reaction path or another, but some things are clear even at this stage. Consider the following example.  [c.16]

The market values and molecular weights of the materials involved are given in Table 2.1. Oxygen is considered to be free at this stage, coming from the  [c.16]

Solution Decisions can be made based on the economic potential of the process (see App. A). At this stage, the best we can do is to define the economic potential (EP) as  [c.17]

If selectivity increases as conversion increases, the initial setting for reactor conversion should he on the order of 95 percent, and that for reversible reactions should be on the order of 95 percent of the equilibrium conversion. If selectivity decreases with increasing conversion, then it is much more difficult to give guidance. An initial setting of 50 percent for the conversion for irreversible reactions or 50 percent of the equilibrium conversion for reversible reactions is as reasonable as can be guessed at this stage. However, these are only initial guesses and will almost certainly be changed later.  [c.27]

Again, it is difficult to select the initial setting of the reactor conversion with systems of reactions in series. A conversion of 50 percent for irreversible reactions or 50 percent of the equilibrium conversion for reversible reactions is as reasonable as can be guessed at this stage.  [c.27]

Multiple reactions. The arguments presented for minimizing reactor volume for single reactions can be used for the primary reaction when dealing with multiple reactions. However, the goal at this stage of the design, when dealing with multiple reactions, is to maximize selectivity rather than to minimize volume for a given conversion.  [c.41]

Homogeneous catalysts. With a homogeneous catalyst, the reaction proceeds entirely in the vapor or liquid phase. The catalyst may modify the reaction mechanism by participation in the reaction but is regenerated in a subsequent step. The catalyst is then free to promote further reaction. An example of such a homogeneous catalytic reaction is the production of acetic anhydride. In the first stage of the process, acetic acid is pyrolyzed to ketene in the gas phase at TOO C  [c.46]

An initial guess for the reactor conversion is very difficult to make. A high conversion increases the concentration of monoethanolamine and increases the rates of the secondary reactions. As we shall see later, a low conversion has the effect of decreasing the reactor capital cost but increasing the capital cost of many other items of equipment in the flowsheet. Thus an initial value of 50 percent conversion is probably as good as a guess as can be made at this stage.  [c.51]

To make an initial guess for the reactor conversion is again diflicult. The series nature of the byproduct reactions suggests that a value of 50 percent is probably as good as csm be suggested at this stage.  [c.52]

Other designs of kilns use static shells rather than rotating shells and rely on mechanical rakes to move solid material through the reactor.  [c.60]

For multiple reactions in which the byproduct is formed in series, the selectivity decreases as conversion increases. In this case, lower conversion than that for single reactions is expected to be appropriate. Again, the best guess at this stage is to set the conversion to 50 percent for irreversible reactions or to 50 percent of the equilibrium conversion for reversible reactions.  [c.64]

It should be emphasized that these recommendations for the initial settings of the reactor conversion will almost certainly change at a later stage, since reactor conversion is an extremely important optimization variable. When dealing with multiple reactions, selectivity is maximized for the chosen conversion. Thus a reactor type, temperature, pressure, and catalyst are chosen to this end. Figure 2.10 summarizes the basic decisions which must be made to maximize selectivity.  [c.64]

Separation of mixtures of condensable and non-condensable components. If a fluid mixture contains both condensable and noncondensable components, then a partial condensation followed by a simple phase separator often can give a food separation. This is essentially a single-stage distillation operation. It is a special case that deserves attention in some detail later.  [c.75]

Even though choices of separators must be made at this stage in the design, it must be borne in mind that the assessment of separation processes ideally should be done in the context of the total system. As is discussed later, separators which use an input of heat to carry out the separation often can be run at effectively zero energy cost if they are appropriately heat integrated with the rest of the process. This includes the three most common types of separators, i.e., distillation columns, evaporators, and dryers. Although they are energy intensive, they also can be energy efficient in terms of the overall process if they are properly heat integrated (see Chaps. 14 and 15).  [c.76]

Partially vaporized feed reverses these effects. For a given separation, the feed conditions can be optimized. No attempt should be made to do this at this stage in the design, since heat integration is likely to change the optimal setting later in the design. It is usually adequate to set the feed to saturated liquid conditions. This tends to equalize the vapor rate below and above the feed.  [c.78]

As with distillation, no attempt should be made to carry out any optimization of liquid flow rate, temperature, or pressure at this stage in the design. The separation in absorption is sometimes enhanced by adding a component to the liquid which reacts with the solute.  [c.84]

Single-stage evaporators tend only to be used when the capacity needed is small. It is more usual to employ multistage systems which recover and reuse the latent heat of the vaporized material. Three  [c.84]

Forward-feed operation is shown in Fig. 3.12a. The fresh feed is added to the first stage and fiows to the next stage in the same direction as the vapor flow. The boiling temperature decreases from stage to stage, and this arrangement is thus used when the  [c.85]

Figure 3.12 Three possible arrangements for a three-stage evaporator. Figure 3.12 Three possible arrangements for a three-stage evaporator.
Backward-feed operation is shown in Fig. 3.126. Here, the fresh feed enters the last and coldest stage and leaves the first stage as concentrated product. This method is used when the concentrated product is highly viscous. The high temperatures in the early stages reduce viscosity and give higher heat transfer coefficients. Because the solutions fiow against the pressure gradient between stages, pumps must be used to transfer solutions between stages.  [c.86]

Parallel-feed operation is illustrated in Fig. 3.12c. Fresh feed is added to each stage, and product is withdrawn from each stage. The vapor from each stage is still used to heat the next stage. This arrangement is used mainly when the feed is almost saturated, particularly when solid crystals are the product.  [c.86]

Many other mixed-feed arrangements are possible which combine the individual advantages of each type of arrangement. Figure 3.13 shows a three-stage evaporator in temperature-enthalpy terms, assuming that inlet and outlet solutions are at saturated conditions  [c.86]

Given these degrees of freedom, how can an initialization be made for the design The most significant degree of freedom is the choice of number of stages. If the evaporator is operated using hot and cold utility, as the number of stages is increased, a tradeoff might be expected, as shown in Fig. 3.14. Here, starting with a single stage, it has a low capital cost but requires a large energy cost. Increasing the stages to two decreases the energy cost in return for a small increase in capital cost, and the total cost decreases. However, as the stages are increased, the increase in capital cost at some point no longer compensates for the corresponding decrease in energy cost, and the total cost increases. Hence there is an optimal number of stages. However, no attempt should be made to carry out this optimization at this point, since the design is almost certain to change significantly when heat integration is considered later (see Chap. 15).  [c.87]

All that can be done is to make a reasonable initial assessment of the number of stages. Having made a decision for the number of stages, the heat flow through the system is temporarily fixed so that the design can proceed. Generally, the maximum temperature in evaporators is set by product decomposition and fouling. Therefore, the highest-pressure stage is operated at a pressure low enough to be below this maximum temperature. The pressure of the lowest-pressure stage is normally chosen to allow heat rejection to cooling water or air cooling. If decomposition and fouling are not a problem, then the stage pressures should be chosen such that the highest-pressure stage is below steam temperature and the lowest-pressure stage above cooling water or air cooling temperature.  [c.87]

The most common alternative to distillation for the separation of low-molecular-weight materials is absorption. Liquid flow rate, temperature, and pressure are important variables to be set, but no attempts should be made to carry out any optimization at this stage.  [c.92]

As with distillation and absorption, when evaporators and dryers are chosen, no attempt should be made to carry out any optimization at this stage in the design.  [c.92]

Stage, H., Fortschritte der Verfahrenstechnik, 2, 306 (1956), 3, 364 (1958) Stage, H., and P. Faldix, Fortschritte der Verfahrenstechnik, 4, 429 (1961), 5, 515 (1962).  [c.12]

To illustrate, UNIQUAC parameters were obtained for the ethanol/cyclohexane system using the extensive isothermal data of Scatchard and Satkiewicz (1964). Figure 2 shows parameters for 5, 35, and 60°C along with the confidence ellipses. These regions indicate that it is possible to choose a single value of 322 appropriate for all temperatures a single value of a2 (e.g. 1300) can be included in all three confidence ellipses, implying that in the range 5-65 C parameter a2 is temperature independent. For 3., however, there is no single value which can intercept all three confidence ellipses. Therefore, parameter a 2 must be represented by a function of temperature as shown in Table 1 where the estimated variance of the fit, a, provides a measure of how well the data are represented. The first line shows results obtained when fitting two UNIQUAC parameters, a 2 21 ii ispendent of temperature. The next two  [c.45]

Figure 4-2. UNIQUAC parameters and their approximate confidence regions for the ethanol-cyclohexane system for three isotherms. Data of Scatchard and Satkiewicz, 1964. Figure 4-2. UNIQUAC parameters and their approximate confidence regions for the ethanol-cyclohexane system for three isotherms. Data of Scatchard and Satkiewicz, 1964.
Then a practical reactor is selected, approaching as nearly as possible the ideal in order that the design can proceed. However, the reactor design cannot be fixed at this stage, since, as we shall see later, it interacts strongly with the rest of the flowsheet. We shall concentrate here on the choice of reactor and not its detailed sizing, which is outside our scope (for the details of sizing, see Denbigh and Turner, Levenspiel, and Rase ).  [c.15]

It is now possible to define the goals for the selection of the reactor at this stage in the design. Unconverted material usually can be separated and recycled later. Because of this, the reactor conversion cannot be fixed finally until the design has progressed much further than just choosing the reactor. As we shall see later, the choice of reactor conversion has a major influence on the rest of the process. Nevertheless, some decisions must be made regarding the reactor for the design to proceed. Thus we must make some guess for the reactor conversion in the knowledge that this is likely to change once more detail is added to the total system.  [c.25]

Unwanted byproducts usually cannot be converted back to useful products or raw materials. The reaction to unwanted byproducts creates both raw materials costs due to the raw materials which are wasted in their formation and environmental costs for their disposal. Thus maximum selectivity is wanted for the chosen reactor conversion. The objectives at this stage can be summarized as follows  [c.25]

The b3TJroduct, DCD, is not required for this project. Hydrogen chloride can be sold to a neighboring plant. Assume at this stage that all separations can be carried out by distillation. The normal boiling points are given in Table 4.1.  [c.102]

At this stage, how great the excess of chlorine should be for Fig. 4.7c to be feasible cannot be specified. Experimental work on the reaction chemistry would be required in order to establish this. However, the size of the excess does not change the basic structure.  [c.104]


See pages that mention the term Saitzew : [c.81]    [c.112]    [c.9]    [c.64]    [c.78]    [c.85]    [c.86]   
Organic syntheses based on name reactions and unnamed reactions (1994) -- [ c.325 ]