Plant fluid catalytic cracking

The MTO process employs a turbulent fluid-bed reactor system and typical conversions exceed 99.9%. The coked catalyst is continuously withdrawn from the reactor and burned in a regenerator. Coke yield and catalyst circulation are an order of magnitude lower than in fluid catalytic cracking (FCC). The MTO process was first scaled up in a 0.64 m /d (4 bbl/d) pilot plant and a successfiil 15.9 m /d (100 bbl/d) demonstration plant was operated in Germany with U.S. and German government support. 

The ethylene feedstock used in most plants is of high purity and contains 200—2000 ppm of ethane as the only significant impurity. Ethane is inert in the reactor and is rejected from the plant in the vent gas for use as fuel. Dilute gas streams, such as treated fluid-catalytic cracking (FCC) off-gas from refineries with ethylene concentrations as low as 10%, have also been used as the ethylene feedstock. The refinery FCC off-gas, which is otherwise used as fuel, can be an attractive source of ethylene even with the added costs of the treatments needed to remove undesirable impurities such as acetylene and higher olefins. Its use for ethylbenzene production, however, is limited by the quantity available. Only large refineries are capable of deUvering sufficient FCC off-gas to support an ethylbenzene—styrene plant of an economical scale. 

The most important undesired metallic impurities are nickel and vanadium, present in porphyrinic structures that originate from plants and are predominantly found in the heavy residues. In addition, iron may be present due to corrosion in storage tanks. These metals deposit on catalysts and give rise to enhanced carbon deposition (nickel in particular). Vanadium has a deleterious effect on the lattice structure of zeolites used in fluid catalytic cracking. A host of other elements may also be present. Hydrodemetallization is strictly speaking not a catalytic process, because the metallic elements remain in the form of sulfides on the catalyst. Decomposition of the porphyrinic structures is a relatively rapid reaction and as a result it occurs mainly in the front end of the catalyst bed, and at the outside of the catalyst particles. 

Table 7 shows the yield distribution of the C4 isomers from different feedstocks with specific processing schemes. The largest yield of butylenes comes from the refineries processing middle distillates and from olefins plants cracking naphtha. The refinery product contains 35 to 65% butanes olefins plants, 3 to 5%. Catalyst type and operating severity determine the selectivity of the C4 isomer distribution (41) in the refinery process stream. Processes that parallel fluid catalytic cracking to produce butylenes and propylene from heavy cmde oil fractions are under development (42). 

The two industrial fluid catalytic cracking units that we consider are of type IV with U-bends. The two units vary in their input parameters which lead to different outputs. The following table contains the plant data for the two commercial FCC units under consideration. 

A model for the riser reactor of commercial fluid catalytic cracking units (FCCU) and pilot plants is developed This model is for real reactors and feedstocks and for commercial FCC catalysts. It is based on hydrodynamic considerations and on the kinetics of cracking and deactivation. The microkinetic model used has five lumps with eight kinetic constants for cracking and two for the catalyst deactivation. These 10 kinetic constants have to be previously determined in laboratory tests for the feedstock-catalyst considered. The model predicts quite well the product distribution at the riser exit. It allows the study of the effect of several operational parameters and of riser revampings. 

Fluid catalytic cracking (FCC) of heavy oil fractions is a well-known process in oil refineries. Numerous books (e.g., 1—3) and publications about the different aspects of this subject are available. This chapter is concerned with the modeling of the transfer line or riser reactor of an FCC unit (FCCU) or of a pilot plant. The riser reactor in FCCUs is a vertical pipe about 1 m in diameter and 10-30 m in height. The hot catalyst coming from the regenerator at about 710 ° C first comes in contact with steam and is fluidized. Then, at a height of some meters above, the catalyst is mixed with the preheated feedstock at about 300 ° C. 

While these techniques have been applied to energy-related processes such as heat-integrated distillation columns and fluid catalytic cracking reactors, there is still extensive research required before the concept of plant design/control is reduced to practice. 

The application of the design methods developed in the previous sections has been demonstrated on a CSTR, a distillation column, fluid catalytic cracking units and a gasoline polymerization plant (20). Here, we will discuss optimizing control of the fluid catalytic cracker. 

One particular advantage of the CDTECH process is the ability of the alkylation reactor to accept a dilute ethylene feed. Because the alkylator operates in a mixed vapor-liquid phase, it is capable of utilizing dilute ethylene feeds, for example, offgas from a fluid catalytic cracking plant or dilute ethylene from a steam... 

Fig. 10.10. Two pioneering fluidized bed reactors (a) the Winkler gas generator (b) the first large-scale pilot plant for fluid catalytic cracking. Reprinted from [82] with permission from Elsevier.

Figure 7. The First Commercial Fluid Catalytic Cracking Plants at Baton Rouge, LA. PCLA 1 at left and at right are two of downflow designs.

Austenitic steels provide excellent corrosion, oxidation, and sulfidation resistance with high creep resistance, toughness, and strength at temperatures greater than 565 °C. Thus they are used in refineries for heater tubes, heater tube supports, and in amine, fluid catalytic cracking (FCC), catalytic hydro-desulfurization (CHD) sulfur, and hydrogen plants. 

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