The coke is stripped with steam in a baffled section at the bottom of the reactor to prevent reaction products, other than coke, from being entrained with the coke leaving the reactor. The coke heater is also a fluidized bed and its primary function is to transfer heat from the gasifier to the reactor.
Coke flows from the coke heater to a third fluidized bed in the gasifier where it is reacted with air and steam to produce a fuel gas product consisting of CO, H 2 , CO 2 , and N 2. This gas flows from the top of the gasifier to the bottom of the heater where it serves to fluidize the heater bed and provide the heat needed in the reactor.
The reactor heat requirement is supplied by recirculating hot coke from the gasifier to the heater. The overall coke inventory in the system is main- tained by withdrawing a stream of purge coke from the heater.
The coke gas leaving the heater is cooled in a waste heat steam generator before passing through external cyclones and a venturi-type wet scrubber. The coke fines collected in the venturi scrubber plus the purge coke from the heater represent the net coke yield and contain essentially all of the metal and ash com- ponents of the reactor feed stock.
After removal of entrained coke fines the coke gas is treated for removal of hydrogen sulfide in a Stretford unit and then used for refinery fuel. In the fluid coking process only enough of the coke is burned to satisfy the heat requirements of the reactor and the feed preheat. The balance of the coke is withdrawn from the burner vessel and is not gasified as it is in a flexicoker.
Therefore, only two fluid beds are used in a fluid coker—a reactor and a burner which replaces the heater. The primary advantage of the Flexicoker Fig. In addition, the coke gas can be used to displace liquid and gaseous hydrocarbon fuels in the refinery process heaters and does not have to be used exclusively in boilers as is the case with fluid coke. Typical yields for many various feeds are available from the licenser.
One set of yields is shown in Table 5. The coke which is gasified in a Flexicoker produces coke gas of the follow- ing approximate composition after H 2 S removal: Table 5. The utility requirements for Fluid Coking are significantly higher than those for delayed coking primarily because of the energy required to circulate the solids between fluid beds.
The air blower in a Flexicoker requires more power than that for a Fluid Coker. The process Licensor should be consulted to determine reasonably accurate utility requirements. Long paraffinic side chains attached to aromatic rings are the primary cause of high pour points and viscosities for paraffinic base residua. Visbreaking is carried out at conditions to optimize the breaking off of these long side chains and their subsequent cracking to shorter molecules with lower viscosities and pour points.
The amount of cracking is limited, however, because if the operation is too severe, the resulting product becomes unstable and forms polymerization products during storage which cause filter plugging and sludge formation. The objective is to reduce the viscosity as much as possible without significantly affecting the fuel stability. The degree of viscosity and pour point reduction is a function of the compo- sition of the residua feed to the visbreaker.
High asphaltene content in the feed reduces the conversion ratio at which a stable fuel can be made [15], which results in smaller changes in the properties. The properties of the cutter stocks used to blend with the visbreaker tars also have an effect on the severity of the visbreaker operation. Aromatic cutter stocks, such as catalytic gas oils, have a favorable effect on fuel stability and permit higher visbreaker conversion levels before reaching fuel stability limitations [17].
The oil fraction is soluble in propane the resin fraction is soluble and the asphaltene fraction insoluble in either pentane, hexane, n- heptane, or octane, depending upon the investigator.
Usually either pentane or n-heptane is used. The solvent selected does have an effect on the amounts and properties of the fractions obtained, but normally little distinction is made in terminology.
Chapter 9 Catalytic Hydrocracking and Hydrocracking and Hydro- processing contains a more detailed discussion of the properties of these frac- tions. Many investigators believe the asphaltenes are not in solution in the oil and resins, but are very small, perhaps molecular size, solids held in suspension by the resins, and there is a definite critical ratio of resins to asphaltenes below which the asphaltenes will start to precipitate. During the cracking phase some of the resins are cracked to lighter hydrocarbons and others are converted to asphaltenes.
Both reactions affect the resin—asphaltene ratio and the resultant stability of the visbreaker tar product and serve to limit the severity of the opera- tion. Cracking of the side chains attached to cycloparaffin and aromatic rings at or close to the ring so the chains are either removed or shortened to methyl or ethyl groups. Cracking of resins to light hydrocarbons primarily olefins and com- pounds which convert to asphaltenes. There are two types of visbreaker operations, coil and furnace cracking and soaker cracking.
As in all cracking processes, the reactions are time—temperature dependent see Table 5. The product yields and properties are similar, but the soaker opera- tion with its lower furnace outlet temperatures has the advantages of lower energy consumption and longer run times before having to shut down to remove coke from the furnace tubes. Run times of 3—6 months are common for furnace vis- breakers and 6—18 months for soaker visbreakers.
This apparent advantage for soaker visbreakers is at least partially balanced by the greater difficulty in clean- ing the soaking drum [2]. Process flow diagrams are shown in Figures 5. The feed is intro- duced into the furnace and heated to the desired temperature.
Typical yields and product properties from visbreaking operations are shown in Tables 5. Many of the properties of the products of visbreaking vary with conversion and the characteristics of the feedstocks.
However, some properties, such as diesel index and octane number, are more closely related to feed qualities; and others, such as density and viscosity of the gas oil, are relatively independent of both conversion and feedstock characteristics [16].
Coking and Thermal Processes 87 Figure 5. The delayed coker material balance is calculated from the equations given in Table 5. The results are tabulated in Table 5. Although at this time the only feed available for the delayed coker is the vacuum tower bottoms stream, other process units in the refinery produce heavy product streams that either can be blended into heavy fuel oil or sent to the delayed coker. The market for heavy fuel oil is limited and, for this example problem, no heavy fuel oil is produced.
These streams are included in the feed to coker in this problem. Sulfur distribution is obtained from Table 5. In actual operations some of the sulfur will be combined as mercaptan molecules R-S-H but for preliminary calculations it is sufficiently accurate to assume that all the sulfur in the gas fraction is combined as H 2 S. For the crude oil in Figures 3. Estimate the coke yields for the crude oil fraction of problem 1 and make a material balance around the delayed coking unit.
Using the information from problem 2, estimate the capital cost of a BPSD delayed coking unit and its utility requirements.
Using Bureau of Mines distillation data from Appendix C, calculate the coke yield and make a material balance from a Torrence Field, California crude oil residuum having an API gravity of Estimate the capital and operating costs for a 10, BPSD delayed coker processing the reduced crude of problem 4. Make a material balance around the delayed coker for the charge rate and the crude oil of problems 6 and 7. Also estimate the utility require- ments. Estimate the capital and operating costs for a 30, BPSD delayed coker processing the assigned reduced crude.
Allan et al. Allan, C. Martinez, C. Eng, and W. Barton, Chem. Akbar and H. Geelen, Hydro. Dymond, Hydro. Elliott, Hydro. Eppard, Petrol.
Refiner, p. Hournac, J. Kuhn, and M. Notarbartolo, Hydro. Hus, Oil Gas J. Lieberman, Oil Gas J. McDonald, Refining Engineer, p. C Sept. Mekler, Refining Engineer, p. C-7 Sept. Mochida, T. Furuno, Y. Korai, and H.
Fujitsu, Oil Gas J. Nagy and L. Antalffy, Oil Gas J. Notarbartolo, C. Menegazzo, and J. Kuhn, Hydro Proc. Remirez, Chem. Rhoe and C. Rose, Hydro. Shea, U. Patent No. Stormont, Oil Gas J. Swain, Oil Gas J. Weisenborn, Jr. Janssen, and T. Hanke, Energy Prog. Stolfa, Hydro. Cooper and W. Reinhold Publishing Corp. Feintuch and K. Meyers, Ed. Originally cracking was accomplished thermally but the catalytic process has almost com- pletely replaced thermal cracking because more gasoline having a higher octane and less heavy fuel oils and light gases are produced [28].
The light gases pro- duced by catalytic cracking contain more olefins than those produced by thermal cracking Table 6. The cracking process produces carbon coke which remains on the catalyst particle and rapidly lowers its activity.
To maintain the catalyst activity at a useful level, it is necessary to regenerate the catalyst by burning off this coke with air. As a result, the catalyst is continuously moved from reactor to regenerator and back to reactor.
The cracking reaction is endothermic and the regeneration reac- tion exothermic. Some units are designed to use the regeneration heat to supply that needed for the reaction and to heat the feed up to reaction temperature. The catalytic-cracking processes in use today can all be classified as either moving-bed or fluidized-bed units. There are several modifications under each of the classes depending upon the designer or builder, but within a class the basic operation is very similar. The Thermafor catalytic cracking process TCC is representative of the moving-bed units and the fluid catalytic cracker FCC of the fluidized-bed units.
The FCC units can be classified as either bed or riser transfer line cracking units depending upon where the major fraction of the cracking reaction occurs. The hot oil feed is contacted with the catalyst in either the feed riser line or the reactor. As the cracking reaction progresses, the catalyst is progressively deactivated by the for- mation of coke on the surface of the catalyst.
The catalyst and hydrocarbon vapors are separated mechanically, and oil remaining on the catalyst is removed by steam stripping before the catalyst enters the regenerator. The oil vapors are taken over- head to a fractionation tower for separation into streams having the desired boil- ing ranges.
The spent catalyst flows into the regenerator and is reactivated by burning off the coke deposits with air. Regenerator temperatures are carefully controlled to prevent catalyst deactivation by overheating and to provide the desired amount of carbon burn-off. The flue gas and catalyst are separated by cyclone separators and electrostatic precipitators. The catalyst in some units is steam-stripped as it leaves the regener- ator to remove adsorbed oxygen before the catalyst is contacted with the oil feed.
The fluidized catalyst is circulated continuously between the reaction zone and the regeneration zone and acts as a vehicle to transfer heat from the regenerator to the oil feed and reactor. Typical FCC unit configurations are shown in Figures 6. One of the most important process differences in FCC units relates to the location and control of the cracking reaction.
Until about , most units were designed with a discrete dense-phase fluidized-catalyst bed in the reactor vessel. The units were operated so most of the cracking occurred in the reactor bed. The extent of cracking was controlled by varying reactor bed depth time and temperature.
Although it was recognized that cracking occurred in the riser feed- ing the reactor because the catalyst activity and temperature were at their highest there, no significant attempt was made to regulate the reaction by controlling riser conditions. After the more reactive zeolite catalysts were adopted by refiner- ies, the amount of cracking occurring in the riser or transfer line increased to levels requiring operational changes in existing units.
As a result, most recently constructed units have been designed to operate with a minimum bed level in Figure 6. Many older units have been modified to maximize and control riser cracking.
Units are also operating with different combinations of feed-riser and dense-bed reactors, including feed-riser followed by dense-bed, feed-riser in parallel with dense-bed, and parallel feed-riser lines one for fresh feed and the other for recycle [31]. The major changes have been to take advantage of im- provements in catalysts and to get more efficient contact of heavy feedstocks with the catalyst particles.
The results have been higher conversion levels with better selectivity higher gasoline yields at given conversion levels by shorter and better controlled reaction times 1—3 sec , closed cyclones, and improved feed distribution systems. These have been summarized by James R. Murphy [26]. Most of the designs are similar to those shown in Figures 6.
The fresh feed and recycle streams are preheated by heat exchangers or a furnace and enter the unit at the base of the feed riser where they are mixed with the hot regenerated catalyst. The heat from the catalyst vaporizes the feed and brings it up to the desired reaction temperature. The mixture of catalyst and hy- drocarbon vapor travels up the riser into the reactors. Kellogg design.
Courtesy of Exxon Research and Engineering. The hydrocarbon vapors are sent to the synthetic crude fractionator for separation into liquid and gaseous products. The catalyst leaving the reactor is called spent catalyst and contains hydro- carbons adsorbed on its internal and external surfaces as well as the coke depos- ited by the cracking. In the regenerator, coke is burned from the catalyst with air. The regenerator temperature and coke burn- off are controlled by varying the air flow rate.
The regenerated catalyst contains from 0. Kellogg design, riser FCC unit. Catalytic Cracking Figure 6. Kellogg resid fluid catalytic cracking RFCC unit.
The regenerator can be designed and operated to burn the coke on the cata- lyst to either a mixture of carbon monoxide and carbon dioxide or completely to carbon dioxide. Newer units are designed and operated to burn the coke to carbon dioxide in the regenerator because they can burn to a much lower residual carbon level on the regenerated catalyst.
This gives a more reactive and selective catalyst in the riser and a better product distribution results at the same equilibrium catalyst activity and conversion level.
For units burning to carbon monoxide, the flue gas leaving the regenerator contains a large quantity of carbon monoxide which is burned to carbon dioxide in a CO furnace waste heat boiler to recover the available fuel energy.
The hot gases can be used to generate steam or to power expansion turbines to compress the regeneration air and generate electric power. This has resulted in units which have a catalyst-oil separator in place of the fluidized-bed reactor to achieve maximum gasoline yields at a given conversion level. These units incorporate a high height-to-diameter ratio lower regenerator section for more efficient single-stage regeneration and an offset side-by-side or stacked vessel design.
UOP-designed units utilize high velocity, low inventory regenera- tors Fig. All units are now designed for both complete combustion to CO 2 and CO combustion control. The large differential value between residual fuel and other catalytic crack- ing feed stocks has caused refiners to blend atmospheric and vacuum tower bot- toms into the FCC feed. Residual feed stocks have orders of magnitude higher metals contents especially nickel and vanadium and greater coke forming poten- tial Ramsbottom and Conradson carbon values than distillate feeds.
These con- taminants reduce catalyst activity, promote coke and hydrogen formation, and decrease gasoline yield. It has been shown that catalyst activity loss due to metals is caused primarily by vanadium deposition, and increased coke and hydrogen formation is due to nickel deposited on the catalyst [6,7].
The high coke laydown creates problems because of the increased coke burning requirement with the resulting increased air or oxygen demand, higher regenerator temperatures, and greater heat removal. Primary reactions are designed as those involving the initial carbon—carbon bond scission and the immediate neutralization of the car- bonium ion [24].
These olefins add a proton from the catalyst to form large carbonium ions which decompose according to the beta rule carbon—carbon bond scission takes place at the carbon in the position beta to the carbonium ions and olefins to form small carbonium ions and olefins. The small carbonium ions propagate the chain reaction by transferring a hydrogen ion from a n-paraffin to form a small paraffin molecule and a new large carbonium ion [11,31].
Step 1: Mild thermal cracking initiation reaction. Even though the basic mechanism is essentially the same, the manner and extent of response to catalytic cracking differs greatly among the various hydro- carbon types. In respect to reaction rates, the effect of the catalyst is more pronounced as the number of carbon atoms in the molecule increases, but the effect is not appreciable until the number of carbon atoms is at least six.
The cracking rate is also influenced by the structure of the molecule, with those containing tertiary carbon atoms cracking most readily, while quaternary carbon atoms are most resistant. Compounds containing both types of carbon atoms tend to neutralize each other on a one-to-one basis. The main reactions are [11]: 1. Carbon—carbon bond scissions 2. Isomerization 3.
Polymerization 4. Saturation, aromatization, and carbon formation Olefin isomerization followed by saturation and aromatization are responsi- ble for the high octane number and lead susceptibility of catalytically cracked gasolines. The higher velocity of hydrogen transfer reactions for branched olefins results in ratios of iso- to normal paraffins higher than the equilibrium ratios of the parent olefins.
In addition, naphthenes act as hydrogen donors in transfer reactions with olefins to yield isoparaffins and aromatics. Dehydrogenation is very extensive for C 9 and larger naphthenes and a high-octane gasoline results. The non-ring liquid products and cracked gases resulting from naphthenic hydro- carbon cracking are more saturated than those resulting from cracking paraffins. The predominant reaction for aromatics with long alkyl chains is the clean splitting off of side chains without breaking the ring.
The carbon—carbon bond ruptured is that adjacent to the ring, and benzene compounds containing alkyl groups can be cracked with nearly quantitative recovery of ben- zene [11]. Catalytic Cracking 6. Most catalysts used in commercial units today are either class 3 or mixtures of classes 2 and 3 catalysts [15]. See Tables 6. The advantages of the zeolite catalysts over the natural and synthetic amorphous catalysts are: 1. Higher activity 2. Higher gasoline yields at a given conversion 3.
Production of gasolines containing a larger percentage of paraffinic and aromatic hydrocarbons 4. Lower coke yield and therefore usually a larger throughput at a given conversion level 5. Increased isobutane production 6. Ability to go to higher conversions per pass without overcracking Table 6. Table 6. Here the adverse effects of carbon deposits on catalyst activity and selectivity are minimized because of the negligible amount of catalyst back- mixing in the riser.
In addition, separate risers can be used for cracking the recycle stream and the fresh feed so that each can be cracked at their own optimum conditions. Amorphous catalysts have higher attrition resis- tance and are less costly than zeolitic catalysts. Lower attrition rates also greatly improve particulate emission rates. Basic nitrogen compounds, iron, nickel, vanadium, and copper in the oil act as poisons to cracking catalysts [1]. The nitrogen reacts with the acid centers on the catalyst and lowers the catalyst activity.
The metals deposit and accumulate on the catalyst and cause a reduction in throughput by increasing coke formation and decreasing the amount of coke burn-off per unit of air by catalyzing coke combustion to CO 2 rather than to CO. Although the deposition of nickel and vanadium reduces catalyst activity by occupying active catalytic sites, the major effects are to promote the formation of gas and coke and reduce the gasoline yield at a given conversion level.
Tolen [39] and others [15] have discussed the effects of nickel and vanadium deposits on equilibrium catalysts Fig. Metals loadings on equilibrium cata- lysts are now as high as 10, ppm and, in , the mean was over ppm of nickel plus vanadium [34].
The effects of nickel can be partially offset by the addition of passivators, such as antimony and barium compounds, to the feed and catalysts containing tin, barium and strontium titanates, and magnesium oxide have been developed which act as metal traps for vanadium [21]. Metals removal processes can also be used to reactivate the catalyst by cycling a slip stream through a metals removal system Demet. This permits the equilibrium catalyst metals concentrations to be controlled at the level at which the fresh catalyst required to maintain activity and selectivity equals catalyst losses [9].
A range of catalysts is available from catalyst manufacturing companies that are compounded to give high gasoline octanes 0. Figure 6. The average microactivity of FCC catalyst in North and South America in was between 67 and 70, while world activity averaged between 65 and Additives are available which increase the percentage of propylene in the butane and lighter gases from the FCC unit.
By operating at higher conversion levels the butylene and propylene production as well as total butane and lighter gas yields can be increased at the expense of naphtha. The naphtha has 1—2 higher blending octane numbers because the additives crack the low octane straight chain and lightly branched paraffins and olefins to propylenes and butylenes [35].
The meso pores are most effective for reducing bottoms yields for aromatic and naphthenic feedstocks and small pores for paraffinic feedstocks. Treating feed to the FCC unit offers the advantages that the sulfur and nitrogen in the gasoline and diesel fuel products are reduced and, by adding hydrogen to the feed, naphtha and LCO yields are increased without lowering the olefins content and octanes of the naphtha fraction [42].
The hydrotreating unit can be operated in several ways: as a hydrodesulfur- ization HDS unit, a mild hydrocracking MHC unit, or a partial-conversion hydrocracking unit. In all cases the product sulfur content has to be less than wppm to produce a refinery gasoline blending pool with less than 50 wppm sulfur and less than 85 wppm to produce a refinery gasoline blending pool of less than 30 wppm.
Higher operating pressures [— psig 95— barg ] are necessary to provide the aromatic ring hydrogenation necessary for partial-conversion hydrocracking Fig.
The partial-conversion hy- drocracking unit also produces distillates with 50 cetane indices CI , less than 50 wppm sulfur, and smoke points of 15—19 mm. Using MHC or partial-conversion hydrocracking instead of hydrocracking offers greater flexibility by increasing the yield of higher-value products such as diesel fuel and jet fuel.
MHC, because of its lower operating pressure and less hydrogenation of aromatic rings, typically produces diesel fuels with cetane indi- ces from 39—42 and the kerosine smoke points well below the 19 mm smoke point required for jet fuels.
Partial-conversion hydrocracking, operating at — Figure 6. For a better understanding of the process, several terms should be defined. Activity: Ability to crack a gas oil to lower boiling fractions. Selectivity: The ratio of the yield of desirable products to the yield of unde- sirable products coke and gas. In a fluidized-bed reactor, the LHSV has little meaning because it is difficult to establish the volume of the bed.
Within the limits of normal operations, increasing 1. Reaction temperature 2. Catalyst activity 4. Contact time results in an increase in conversion, while a decrease in space velocity increases conversion. In many FCC units, conversion and capacity are limited by the regenerator coke burning ability. This limitation can be due to either air compression limitations or to the afterburning temperatures in the last stage regenerator cyclones.
In either case FCC units are generally operated at the maximum practical regenerator temperature with the reactor tem- perature and throughput ratio selected to minimize the secondary cracking of gasoline to gas and coke.
With the trend to heavier feedstocks, the carbon forming potential of catalytic cracker feeds is increasing, and some units limited in carbon burning ability because of limited blower capacity are adding oxygen to the air to the regenerator to overcome this limitation. In fluidized-bed units, the reactor pressure is generally limited to 15 to 20 psig by the design of the unit and is therefore not widely used as an operating variable.
Increasing pressure increases coke yield and the degree of saturation of Table 6. It has little effect on the conver- sion. Typical operations of these units are given in Table 6.
Temperature control in the regenerator is easier if the carbon on the catalyst is burned to carbon dioxide rather than carbon monoxide but much more heat is evolved and regenerator temperature limits many be exceeded. A better yield structure with lower coke laydown and higher gasoline yield is obtained at a given conversion level when burning to carbon dioxide to obtain a lower residual carbon on the catalyst. Many catalytic crackers include waste heat boil- ers which recover the sensible heat by steam generation and others use power recovery turbines to generate electric power or compress the air used in the cata- lytic cracker regenerator.
Some refineries recover the heat of combustion of the carbon monoxide in the flue gas by installing CO-burning waste heat boilers in place of those utilizing only the sensible heat of the gases. An even higher rate of energy recovery can be achieved by using a power recovery turbine prior to the CO or waste heat boiler, although when regenerator pressures are less than 15 psig, power recovery turbines usually are not economic.
The correlations are very useful for esti- mating typical yields for preliminary studies and to determine yield trends when changes are made in conversion levels. The yield structure is very dependent upon catalyst type. Allen Pat B. Anderson Richard E.
Blundell Bartholomew Craig G. Bayer Ronald Beers V. Coogan Michael D. Crutchley Lee David S. Feder Kenneth L. Fideler Elizabeth F. Finando Donna Finocchiaro Maurice A. Harrington Jr. Stuecklen Kate M. Linker Maureen Lisa A. Nardozzi Charlie Nesheim John L. Newmark Amy Nigrini Mark J. Steiner Jr. John E. Stevens Dannelle D. Stilton Geronimo Stine R. Thomas Cleary Thomas H. Written by experts with both academic and professional experience in refinery operation, design, and evaluation, Petroleum Refining Technology and Economics, Fifth Edition is an essential textbook for students and a vital resource for engineers.
This latest edition of a bestselling text provides updated data and addresses changes in refinery feedstock, product distribution, and processing requirements resulting from federal and state legislation. It also contains end-of-chapter problems and an ongoing case study. This book is targeted to benefit the diploma in engineering students. Degree in engineering students B.
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