Reactor capacity

Initially, all of the SBR polymer known as GR-S produced during World War II was by the batch process. Later, it was thought that a higher volume of polymer would be needed for the war effort. The answer was found in switching from batchwise to continuous production. This was demonstrated in 1944 at the Houston, Texas, synthetic mbber plant operated by The Goodyear Tire Rubber Company. One line, consisting of 12 reactors, was lined up in a continuous mode, producing GR-S that was mote consistent than the batch-produced polymer (25). In addition to increased productivity, improved operation of the recovery of monomers resulted because of increased (20%) reactor capacity as well as consistent operation instead of up and down, as by batchwise polymerisation. 

REACTCP Reactor capacity limit PCF Product C to fractionation section... 

Flows (columns) FI Cl through F4C3 each create units of REACTL the total flow to the reactor and use up units REACTCP reactor capacity. Flows FlCl through F4C3 create yields of reactor products RPA through RPE. Pure component C is added from an outside source into tank (row) RPC by flow (column) OUTSIDE. 

Only 10 firms account for 75% of agrochemicals sales, while the 15 largest drug companies have a market share of only 33% (Stinson, 1995). About 85% of fine chemicals are manufactured by companies of the triad the United States (28%), Western Europe (39%), and Japan (17%). Italy, with 4.0 million litres reactor capacity and 71 manufacturers, topped the European fine chemicals industry (Layman, 1993). Recently India, China, and Eastern-Central European countries have gained a significant proportion of the market, as a result of the lower direct labour costs and the more relaxed environmental and safety standards. It is fair to state that the high quality of chemists in these countries has also contributed to this development. In 1993, the cost of producing fine chemicals in India was 12% below that in Europe (Layman, 1993). 

Generally, in MPPs one centrifuge or pressure/vacuum filter, one drier, and one fractionation system of reasonable capacity are installed for every two or three reactors/crystallizers. On average, about 2 to 5 m solvent storage capacity and the same amount for intermediates are needed per 1 m of reactor capacity. 

Figure 7.3-1. Total investment costs versus reactor capacity.

Suppose Table 1-1 represents the yield obtained vs. time for each reactor cycle. If the reactor cycle is 8 hours and produces 15,000 lb of product per batch, then if the cycle time were cut to 5 hours the yield would be 13,250 lb per batch. The rates of production would be 1,875 lb / hr for the former and 2,650 lb / hr for the latter. For a plant operated 8,000 hours per year this would give a production rate of 15,000,000 lb / yr for the former and 21,206,000 lb / yr for the latter. A change of this sort would necessitate no increase in reactor capacity, but it would require changes in the recovery and recycle systems other than those solely due to the increase in capacity. 

Operating costs associated with advanced oxidation systems are a function of capacity. One Ultrox unit installed in New York for groundwater treatment of trichloroethylene (TCE) and toluene had a capital cost of approximately 1 million with a 3900-gal reactor capacity and a 250-gal/min flow through capacity. Operating costs for the unit are approximately 1.57 per 1000 gal treated at the flow rate of 250 gal/min (D124629, pp. 736-737). 

This reduction in capital costs per ton of LDPE was made possible by the increase in reactor capacity. Fig. 8.2-2 shows the development of name-plate capacity of LDPE reactors. A steep increase occurred since 1980. Whereas in this year the maximum stream-size was in the range of 80,000 - 120,000 t/a, the maximum capacity of today s reactors is 300,000 t/a. 

When reactor capacity is limited by heat removal, an often-recommended control structure is to run with maximum coolant flow and manipulate feed flowrate to control reactor temperature (Tr F0 control). This control scheme has the potential to achieve the highest possible production rate. However, if the feed temperature is lower than the reactor temperature, the transfer function between temperature and feed flowrate contains a positive zero, which degrades dynamic performance, as we demonstrate quantitatively in this section. The choice of a control structure for this process presents an example of the often encountered conflict between steady-state economics and dynamic controllability. 

Use the kinetic model in Appendix 13.1 to design a CSTR for the production of polystyrene. The entering feed is pure styrene. It is desired to produce 50% by weight of polystyrene with a number average molecular weight of 85,000. The feed flow rate is 25,000 kg/h. Determine the required operating temperature and reactor capacity (in mass units). 

Evidently, changes in the reactor size impact on the above findings allowing an increase in the reactor holdup leads to an increase in the single-pass conversion and reduces the flow rate of the material recycle stream. While plant configurations with low reactor capacity are preferred in processes featuring multiple reactions with valuable intermediate products (Luyben 1993b), the optimal sizes... 

Probably, for most slurry reactor applications, information on the value of the product kLa is sufficient for design purposes. In some cases, however, information on the individual parameters a and/or ki, can be useful. For instance, the reactor capacity will depend on a, rather than on the product k a, if the reaction is so fast that all conversion takes place within the stagnant film (film theory) around the gas bubbles. For first-order conversion kinetics in the porous catalyst particles this will occur for... 

Commercial plants The process is used in 20 reactors at 15 sites with annual single reactor capacities up to 320,000 mtons of EDC, alone as HCI-consuming plant or as part of the balanced VCM process. In some cases, it has replaced other oxychlorination technologies from different licensors by replacing existing reactors or existing units. Two new oxychlorination trains were successfully commissioned in September 2004 one oxychlorination unit is under design. 

In this regime the conversion rate (hence, reactor capacity) can be enhanced by increasing the mass-transfer rate (e.g., by increasing the fluid velocity in the reactor or decreasing the average particle diameter). 

Reactor capacity per unit volume appears to depend on four resistances in series the gas-phase transfer resistance, two liquid-phase transfer resistances, and the kinetic resistance. The highest resistance limits the capacity of the reactor. The four resistances have the unit of time and each one individually represents the time constant of the particular process under study. For example, 1 lkjigl is the time constant for the transfer of A from the bulk of the gas through the gas film to the gas-liquid interface. The same holds for the three other resistances. For a first-order reaction in a batch reactor, for example, the concentration after a certain time is given by C/C0 = exp(-r/r), in which r = 1/ A is the reaction time constant. For processes in series the individual time constants can be added to find the overall time constant of the total process. 

Substrate Concentration Effects on the Reaction Rate, Enzyme Stability, Substrate Conversion, and Reactor Capacity... 

Figure 17.6a shows the effect of the substrate concentration on the regime conversion and the UF-reactor capacity. A 50% substrate conversion was the maximum attained with 2—4mM of substrate, while increasing the substrate concentration to over 4mM led to a decreased conversion. The reactor capacity increased with the substrate concentration and reached a plateau. Since the goal might be to increase the product concentration in the reactor permeate, the enzyme load in the reactor was doubled and, as expected, the steady-state conversion increased from 50 to 62.5% (run fed with 4mM of substrate). However, the reactor capacity dropped from 2.17 to 1.24 a,molmgD cwh . ... 

The space velocity effect was investigated and data analysis allowed evaluation of the steady-state conversion of the reactor loaded with 10 mg and a substrate feed of 4mM in buffer (50mM of sodium phosphate buffer at pH 7.0) solution. These data are illustrated in Figure 17.6b. An increase in T gives a higher product conversion but the reactor capacity is lower. 

Collectively, the experiments clarify the optimal mean residence time (see Figure 17.6b) and cell loading (see Figure 17.7). The optima occur whilst the conversions attained in the CSMR are 30 and 50%, respectively. Varying these operational parameters, the CSMR could produce well-concentrated streams at a still acceptable reactor capacity. Due to the complex effect of the substrate concentration on the reaction rate, the highest conversion is attained at a reactor capacity approximately half that of the maximum (see Figure 17.6a). 

X-Mean residence time (h) Figure 17.6 UF-membrane bioreactor ioaded with lOmgDcw, and run at 10 C. The reactor was fed with benzonitriie buffered (50 mM sodium phosphate buffer, pH 7.0) soiution. (a) % Conversion at steady state ( ) and reactor capacity (o) as a function of... 

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