CHAP process

The CHAP process is based on the hydrolysis of lignocellulosic material by concentrated hydrochloric acid at low temperature and subsequent sugar fermentation. 

Since process design starts with the reactor, the first decisions are those which lead to the choice of reactor. These decisions are among the most important in the whole design. Good reactor performance is of paramount importance in determining the economic viability of the overall design and fundamentally important to the environmental impact of the process. In addition to the desired products, reactors produce unwanted byproducts. These unwanted byproducts create environmental problems. As we shall discuss later in Chap. 10, the best solution to environmental problems is not elaborate treatment methods but not to produce waste in the first place. 

If the total heat consumed is from an external utility (e.g., mains steam), then a high efficiency is desirable, even perhaps at the expense of a high capital cost. However, if the heat consumed is by recovery from elsewhere in the process, as is discussed in Chap. 15, then comparison on the basis of dryer efficiency becomes less meaningful. 

Achieving complete conversion of FEED to PRODUCT in the reactor usually requires an extremely long residence time, which is normally uneconomic (at least in continuous processes). Thus, if there is no byproduct formation, the initial reactor conversion is set to be around 95 percent, as discussed in Chap. 2. The reactor effluent thus contains unreacted FEED and PRODUCT (Fig. 4.1a). 

Clearly, the time chart shown in Fig. 4.14 indicates that individual items of equipment have a poor utilization i.e., they are in use for only a small fraction of the batch cycle time. To improve the equipment utilization, overlap batches as shown in the time-event chart in Fig. 4.15. Here, more than one batch, at difierent processing stages, resides in the process at any given time. Clearly, it is not possible to recycle directly from the separators to the reactor, since the reactor is fed at a time different from that at which the separation is carried out. A storage tank is needed to hold the recycle material. This material is then used to provide part of the feed for the next batch. The final flowsheet for batch operation is shown in Fig. 4.16. Equipment utilization might be improved further by various methods which are considered in Chap. 8 when economic tradeoffs are discussed. 

Whether heat integration is restricted to the separation system or allowed with the rest of the process, integration always benefits from colder reboiler streams and hotter condenser streams. This point is dealt with in more general terms in Chap. 12. In addition, when column pressures are allowed to vary, columns with smaller temperature differences are easier to integrate, since smaller changes in pressure are required to achieve suitable integration. This second point is explained in more detail in Chap. 14. 

Consider again the simple process shown in Fig. 4.4d in which FEED is reacted to PRODUCT. If the process usbs a distillation column as separator, there is a tradeofi" between refiux ratio and the number of plates if the feed and products to the distillation column are fixed, as discussed in Chap. 3 (Fig. 3.7). This, of course, assumes that the reboiler and/or condenser are not heat integrated. If the reboiler and/or condenser are heat integrated, the, tradeoff is quite different from that shown in Fig. 3.7, but we shall return to this point later in Chap. 14. The important thing to note for now is that if the reboiler and condenser are using external utilities, then the tradeoff between reflux ratio and the number of plates does not affect other operations in the flowsheet. It is a local tradeoff. 

Product removal during reaction. Separation of the product before completion of the reaction can force a higher conversion, as discussed in Chap. 2. Figure 2.4 showed how this is done in sulfuric acid processes. Sometimes the product (or one of the products) can be removed continuously from the reactor as the reaction progresses, e.g., by allowing it to vaporize from a liquid phase reactor. 

In Chap. 10, modification of the process for reducing process waste was considered in detail. It also was concluded that to minimize utility waste, the single most effective measure would be improved heat recovery. The energy-targeting methods presented in Chaps. 6 and 7 maximize heat recovery for a given set of process conditions. However, the process conditions can be changed to improve the heat recovery further. 

Having to readjust the capital/energy tradeoff after every process change would be a real problem if it were not for the existence of the total cost targeting procedures discussed in Chap. 7. 

The preceding appropriate placement arguments assume that the process has the capacity to accept or give up the reactor heat duties at the given reactor temperature. A quantitative tool is needed to assess the capacity of the background process. For this purpose, the grand composite curve can be used and the reactor profile treated as if it was a utility, as explained in Chap. 6. 

Distillation capital costs. The classic optimization in distillation is to tradeoff capital cost of the column against energy cost for the distillation, as shown in Fig. 3.7. This wpuld be carried out with distillation columns operating on utilities and not integrated with the rest of the process. Typically, the optimal ratio of actual to minimum reflux ratio lies in the range 1.05 to 1.1. Practical considerations often prevent a ratio of less than 1.1 being used, as discussed in Chap. 3. 

It was noted earlier that dryers are quite difierent in character from both distillation and evaporation. However, heat is still taken in at a high temperature to be rejected in the dryer exhaust. The appropriate placement principle as applied to distillation columns and evaporators also applies to dryers. The plus/minus principle from Chap. 12 provides a general tool that can be used to understand the integration of dryers in the overall process context. If the designer has the freedom to manipulate drying temperature and gas flow rates, then these can be changed in accordance with the plus/minus principle in order to reduce overall utility costs. 

In the volume elements describing individual subchains, the x, y, and z dimensions will be different, so Eq. (3.32) must be averaged over all possible values to obtain the average entropy change per subchain. This process is also easily accomplished by using a result from Chap. 1. Equation (1.62) gives the mean-square end-to-end distance of a subchain as n, 1q, and this quantity can also be written as x + y + z therefore... 

Equivalent mechanical behavior can be achieved by either time (or frequency) or temperature manipulation. As noted in Sec. 3.2, results measured at different temperatures can be reduced to a common temperature to describe response over a wide range of times. We shall consider data reduced to a common temperature in this chapter and discuss the reduction process in Chap. 4. 

The assumption that k values are constant over the entire duration of the reaction breaks down for termination reactions in bulk polymerizations. Here, as in Sec. 5.2, we can consider the termination process—whether by combination or disproportionation to depend on the rates at which polymer molecules can diffuse into (characterized by kj) or out of (characterized by k ) the same solvent cage and the rate at which chemical reaction between them (characterized by kj.) occurs in that cage. In Chap. 5 we saw that two limiting cases of Eq. (5.8) could be readily identified ... 

When we discussed random walk statistics in Chap. 1, we used n to represent the number of steps in the process and then identified this quantity as the number of repeat units in the polymer chain. We continue to reserve n as the symbol for the degree of polymerization, so the number of diffusion steps is represented by V in this section. 

In Chap. 8 we saw how the equilibrium osmotic pressure of a solution is related to AG for the mixing process whereby the solution is formed. Any difference in the concentration of the solution involves a change in AG j, ... 

A. T. Kuhn, Industrial Electrochemical Processes, Elsevier Publishing Co., Amsterdam, The Netherlands, 1971, Chap. 3. 

Bolles, chap. 14 in Smith, Design of Equilihiium Stage Processes, McGraw-Hill, New York, 1963. To convert feet to meters, multiply hy 0.3048 to convert gallons per minute to decimeters per second (hters per second), multiply hy 0.06309 and to convert gallons per minute to ciihic meters per second, multiply hy 6.309 X 10" ... 

For bubble caps, Ki is the drop through the slots and Ko is the drop through the riser, reversal, and annular areas. Equations for evaluating these terms for various bubble-cap designs are given by BoUes (in chap. 14 of Smith, Equilibrium Stage Processes, McGraw-HiU, New York, 1963), or may be found in previous editions of this handbook. 

FIG. 16-57 Sanmatsii Kogyo chromatographic process. (Reptinted with permission of Wiley. Reference Keller, Anderson, and Yon, Chap. 12 in Rousseau, Handbook of Separation Process Technology, John Wiley [Pg.1556]

Cracking was caused by stress-corrosion cracking (see Chap. 9, Stress-Corrosion Cracking ) involving hydrogen sulfide and/or moist sulfur dioxide. The sulfur entered the cooling water stream through process leaks, which were repaired. 

Thiobacillus thiooxidans is an aerobic organism that oxidizes various sulfur-containing compounds to form sulfuric acid. These bacteria are sometimes found near the tops of tubercles (see Chap. 3, Tubercu-lation ). There is a symbiotic relationship between Thiobacillus and sulfate reducers Thiobacillus oxidizes sulfide to sulfate, whereas the sulfate reducers convert sulfide to sulfate. It is unclear to what extent Thiobacillus directly influences corrosion processes inside tubercles. It is more likely that they indirectly increase corrosion by accelerating sulfate-reducer activity deep in the tubercles. 

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