RU2091433C1 - Method of processing lubricating-oil fraction - Google Patents

Abstract

FIELD: petroleum processing. SUBSTANCE: fraction is submitted to catalytic cracking in reactor with ascending zeolite-containing catalyst bed. As lift gas, petroleum processing gas containing at least 10 vol.% hydrogen and less than 10 vol.% C₃-hydrocarbons is used. Contact time of lift gas with suspended catalyst particles is 0.01 to 0.5 s and content of coke on catalyst is no more than 0.20 wt %. EFFECT: enhanced efficiency of process. 6 dwg, 1 tbl

Description

 The invention relates to a method for processing a petroleum oil fraction and can be used in the oil refining industry.

There is a known method [1] in which C ₁ -C ₅ hydrocarbons are mixed with a hot regenerated catalyst that is returned to the cracking stage in the ascending layer until the oil is fed into the reactor with the ascending catalyst layer. Light hydrocarbons are supplied at a rate of about 35.6-89.0 m ³ / m ^{3 of} raw materials.

 According to the method [2], light hydrocarbons are introduced into the cube of the cracking zone with an ascending catalyst bed. Gas oil is higher than the riser, then hydrocarbon is injected even higher than the riser. The suspended catalyst moves upward through these injection sites and is lifted by hydrocarbon vapors and loaded light hydrocarbons.

 A known method [3] in which a heavy hydrocarbon is contacted with a catalyst in the presence of substantially pure hydrogen in an upward stream. Hydrogen is produced outside the cracking unit. It is indicated that hydrogen with an oil feed is used to reduce the formation of coke and to reduce the formation of unsaturated products.

 In the process according to the method [4], the catalyst is contaminated with nickel and vanadium when the catalyst is contacted with hydrogen saturated with water before the catalyst is introduced into the cracking zone in the bottom layer.

 Known methods [5, 6] in which a metal-contaminated cracking catalyst is removed from the cracking zone in the ascending layer to a regeneration zone where coke is burned. The regenerated catalyst is then treated with hydrogen, carbon monoxide or a mixture of hydrogen with carbon monoxide in the reduction zone before it is introduced into the cracking zone.

 According to the method [7], the regenerated catalyst is sent to the reduction zone, where it is contacted with hydrogen. The reduced catalyst and unused hydrogen are then fed to the cracking zone in the bottom layer.

 A known method [8] in which the regenerated metal-contaminated catalyst is treated with a hydrocarbon gas containing three carbon atoms or less, or a mixture thereof, to achieve complete reduction of the polluting metals that are carbonized in an immersion elbow located between the regenerator and the cracking reaction zone in the ascending layer. This immersion elbow is used to transfer the catalyst from the regenerator to the cracking zone in the bottom layer.

 In a method similar to that described above [9], the used reducing gas is a mixture of one, two, three carbon atoms for passivation of metals in a metallic state before their carbonization.

 Method [10] describes an upstream reactor system for cracking an oil feed, such as heavy hydrocarbon residues, distillation residues of crude oil distillation under vacuum or atmospheric pressure, pecks, asphalt, or a mixture in which hot coke particles are first dispersed in a fluidizing gas such as steam, light hydrocarbons, inert gas, or a mixture thereof.

The method [11] describes the extraction of a hydrocarbon product stream boiling below about C ₆ or C ₅ hydrocarbons, which is used in an amount sufficient to form a suspension with particles of a zeolite catalyst in the lower part of the reaction zone with an ascending layer before loading the oil supply to contact them .

A known method [12] which describes the use of a C ₃ -C ₄ mixture of hydrocarbon gases in the bottom of the reactor with an ascending layer for the formation of a rising suspension of particles of a zeolite catalyst before contacting with an oil feed.

 Method [13] describes a cracking process in a fluidized bed of solid particles, in which a gaseous cracking product is used, including hydrogen, methane, unsaturated and saturated gaseous hydrocarbons.

 The method [14] relates to a process for the conversion of carbometallic oil using hydrogen in an ascending bed reactor and a catalyst containing heavy metals. This patent describes the introduction of hydrogen into the reaction zone with an ascending catalyst bed to reduce the formation of conjugated diolefins. It is postulated that the concentration of the diolefin decreases and coke production on the catalyst containing metals also decreases. The patent describes the use of hydrogen gas mixed with a hydrocarbon feed, the formation of gas streams containing 60-80 vol. or more hydrogen.

 Closer to the invention is a method of processing the oil oil fraction by catalytic cracking in the presence of a cracking catalyst containing crystalline zeolite in the cracking zone of the reactor with an ascending catalyst bed [15]. In this case, a hydrocarbon lift gas is used to form an ascending layer of a suspension of catalyst particles. The resulting suspension is then contacted with the feed at a hydrocarbon conversion temperature to obtain the desired products.

 Moreover, the selectivity of the process in relation to the target gasoline and light cyclic oil is insufficient.

 The aim of the invention is to increase the selectivity for the formation of gasoline and light cyclic oil.

The goal is achieved by the described method of processing the oil oil fraction by catalytic cracking in the presence of a cracking catalyst containing crystalline zeolite in the cracking zone of the reactor with an upward catalyst bed using oil refining processes containing at least 10 vol. hydrogen and less than 10 vol. C ₃₊ hydrocarbons for the formation of an ascending layer of a suspension of catalyst particles and subsequent contacting of the suspension with the feedstock at a hydrocarbon conversion temperature to obtain the desired products with the process carried out at a contact time between the elevator gas and suspended catalyst particles of 0.01-0.5 s , and maintaining the coke content on the catalyst with a lift gas of not more than 0.20 wt.

 Distinctive features of the method are the use of an elevator gas of the above composition, as well as the process at a contact time between the elevator gas and suspended catalyst particles of 0.01-0.5 s, while maintaining the coke content on the catalyst with the elevator gas not more than 0.20 wt.

 Characteristics of the catalyst.

 The catalyst used in the hydrocarbon catalytic cracking operation of the present invention can be essentially any liquefiable crystalline cracking zeolite catalyst known in the art containing rare earth and / or hydrogen ions in the crystalline structure of the zeolite. Zeolite is dispersed in a silicon oxide-clay matrix material, which may have some cracking activity. Thus, the matrix can be selected among mixtures of silicon oxide-alumina, silicon oxide-zirconium oxide or silicon oxide-chromium oxide, which are promoted by one or more metal additives, which are effective in the passivation of accumulated metal contaminants. Some additional materials that can be used include rare earth metals, providing an excess of lanthanum and compounds of antimony and titanium. The cracking catalyst used in the method of the present invention may contain an active crystalline zeolite component in an amount of less than about 40 wt. preferably in the range of 5-20 wt.

 A particularly preferred class of catalysts includes catalysts that are capable of hydrogen activation and have a porous structure into which the molecules of the feedstock can enter by adsorption and / or by contact with the active sites of the catalyst inside or near the pores. Various types of catalysts are available within this classification, including, for example, layered silicates.

 The zeolite-containing catalysts used in the present invention can include any zeolite, both natural and semi-synthetic or synthetic, alone or mixed with other materials that do not substantially impair the suitability of the catalyst, provided that the resulting catalyst has activity and structure then described below. For example, if the catalyst is a mixture, it may include a zeolite component dispersed in a porous refractory inorganic oxide support. In this case, the catalyst may, for example, contain about 1-60 wt. more preferably 1-40% and more typically about 5-40 wt. zeolite dispersed in carriers, counting on the total weight of the catalyst (a base that does not contain water).

Catalysts having a particle size in the range of 5 • 10 ^-6 160 • 10 ^-6 m (about 5-160 microns), more preferably 40 • 10 ^-6 120 • 10 ^-6 m (about 40-120 microns), and containing proportionally large number of particles in the range 40 • 10 ^-6 80 • 10 ^-6 m (40-80 microns).

 It is preferable to use a catalyst that initially has a relatively high level of cracking activity and selectivity and provides high levels of conversion and productivity with a short residence time. The conversion ability of the catalyst can be expressed in terms of the conversion obtained during the actual operation or in a standard test for checking catalytic activity.

 For example, it is preferable to use catalysts which, during a long operation in the process, are active enough to maintain a conversion level of at least about 50 wt. or more, preferably at least about 60 wt. In this regard, the conversion is expressed in liquid volume percent, counting on fresh feed.

 A preferred catalyst can also be defined as a catalyst that, when not in use or in equilibrium, shows a certain activity, expressed as a volume percent, corresponding to the micro-activity test (MAT).

 An acceptable level of equilibrium MAT activity of the catalyst, which was used in the method of the present invention, is more than 20 vol. preferably at least about 40 vol. or more preferably about 60 vol.

 Addition of catalyst.

 Typically, the mass ratio of catalyst to fresh feed (feed that has not previously been exposed to cracking catalyst under cracking conditions) used in the present invention is 3-18. The preferred ratio may be about 4-12 depending on the supply trend towards coke formation.

 The catalyst may be added continuously or periodically, for example, to compensate for the usual loss of catalyst from the system. In addition, the addition of the catalyst can be carried out in combination with the removal of the catalyst, for example, so as to maintain or increase the average level of activity of the catalyst in the installation or to maintain a constant amount of metal on the catalyst.

For example, the rate at which fresh catalyst is added to the plant is 0.285 8.55 kg / m ³ feed or more depending on the metal content in the feed and the level of metal that can remain on the equilibrium catalyst. If, on the other hand, an equilibrium catalyst is used, a substitution rate as high as about 14.25 kg / m ^{3 of} feed or more can be used. When the conditions are such that the conditions in the installation tend to promote rapid deactivation, higher administration rates may be used than indicated above, but in opposite circumstances lower administration rates may be applied.

Metal-on-catalyst
The invention can be practiced with a catalyst carrying accumulated heavy metals, which until now have been considered completely unacceptable in traditional processes of liquid catalytic cracking of vacuum gas oil. The use of a catalyst carrying accumulated heavy metals in the range of 1000 to 20,000 ppm Ni + Va on average is contemplated. Accumulation can also be in the range of 4000-50000 ppm or more.

 Heavy metals can be accumulated, including vanadium nickel, iron increment (which is additional iron accumulated with earlier use of the catalyst), and copper. An equilibrium catalyst can be used from another installation, such as a catalytic cracking unit, which was used to crack a vacuum gas oil having a carbon pyrolysis residue of less than 1 and containing less than about 4 ppm nickel equivalents of heavy metals.

Catalyst promoters
The catalyst composition may also include one or more combustion promoters that are useful in the subsequent catalyst regeneration step. To restore the activity of the catalyst, coke is burned at the regeneration stage, in which the coke is converted to combustion gases, including carbon monoxide and / or carbon dioxide. Various substances are known, for example, Pt, Pd, rare earths, which when introduced into cracking catalysts in small amounts (or added to the feed) tend to promote the conversion of coke to carbon monoxide and / or carbon dioxide. Carbon monoxide combustion promoters tend to lower the temperature at which a predetermined degree of coke removal can be achieved, therefore, the potential for thermal deactivation of the catalyst is reduced.

 Such promoters are usually used in effective amounts ranging from trace amounts to 10% 20 wt. catalyst.

Additional materials
The amount of additional materials that may be present in the feed may vary as described, but the amount indicated should preferably be sufficient for the heat balance of the process. These materials can, for example, be introduced into the reaction zone in a mass ratio with respect to a feed of 0.02-0.4, more preferably 0.03-0.3, even more preferably 0.05-0.25.

 When introducing liquid water recycled from the regeneration zone to the reaction zone as additional material or already mixed with the feed or separately, the preferred option is the presence of hydrogen sulfide dissolved in it within the above limits, taking into account the total amount of feed. Alternatively, approximately 50-5000 ppm of hydrogen sulfide may be dissolved in recycled liquid water. Gaseous hydrogen sulfide can also be added as an additional material in the above ranges of mass ratios instead of hydrogen sulfide dissolved in recycled liquid water.

 In the method of the present invention, ballistic separation of the catalyst and the vapors at the lower end of the progressive flow-through progressive bed reactor is used.

 However, depending on whether there is adhesion between the catalyst and the hydrocarbon vapor in the upflow reactor, the residence time of the catalyst in the upflow reactor may coincide with the vapor residence time or not. Thus, the ratio of the average residence time of the catalyst in the reactor to the residence time of the vapor in the reactor can be in the range of 1-5, more preferably 1-4, even more preferably 1.1-3, the preferred range is 1.2-2.

 It is considered advantageous if the residence time of the vapor in the riser and the contact time of the vapor-catalyst in this reactor are substantially the same for at least about 80% with respect to the length of the riser.

Catalyst regeneration
The regeneration of the catalyst can be carried out at a temperature of 593-871 ^o C, measured at the outlet of the catalyst regenerator. The temperature may be 649-816 ^o C, more preferably 677-774 ^o C, optimally 704-746 ^o C or about 760 ^o C.

To minimize the temperature of regeneration and the requirements for regeneration performance, it is desirable to use such values of time, temperature and atmosphere in the steamer that are sufficient to reduce the potentially volatile hydrocarbon material carried by the steaming catalyst to 10% or less of the mass introduced into the regenerator. Such steaming may, for example, include reheating the catalyst, extensive steam stripping, the use of gases having a temperature significantly higher than usual in FCC / VGO operations, such as in the case of gas leaving the regenerator, as well as other gas streams from refineries, such as hydroprocessing waste gas (containing H ₂ S), hydrogen, and others. The steamer can operate at temperatures above about 482 ^o C. Steaming operations, in which the temperature of the spent catalyst is increased to higher temperatures, are also included in the scope of the present invention.

 To maintain the desired activity of the zeolite catalyst, it is desirable to regenerate the catalyst under conditions of time, temperature and atmosphere sufficient to reduce the mass of carbon remaining on the catalyst to 0.05 wt. or less, regardless of whether the catalyst carries a large amount of accumulated heavy metals or not. It should be understood that the term coke includes any residual unevaporated feed or carbonaceous material present on the catalyst after evaporation.

 Significant levels of conversion achieved in the method of the present invention are the result of relatively large yields of coke, such as, for example, 4-1 wt. counting on fresh serving, more usually 6-14 wt. and even more often about 6-12 wt.

According to a preferred embodiment of the present invention, the regeneration process as a whole can be carried out at low levels of coke on a regenerated catalyst with oxygen supplied to one or more stages of regeneration in the stoichiometric amount required to burn all the carbon in the coke to CO and / or CO ₂ , and burning all of the hydrogen in coke to H ₂ O, and burning all of the sulfur in coke to SO ₂ . If the coke includes other combustibles, the aforementioned stoichiometric amount can be adjusted to include the amount of oxygen required to burn them.

Multistage regeneration offers a technique of combined oxygen-deficient regeneration with control of the molar ratio of CO: CO ₂ and still provides values at which the coke on the catalyst is reduced preferably to 0.05 wt. or lower. So, about 65-80 wt. coke on the catalyst is removed in the first stage of regeneration, in which the molar ratio of CO-CO ₂ is controlled.

 The remaining coke can be removed in a subsequent regeneration step in which more oxygen is present.

A particularly preferred embodiment of the present invention is a two-stage catalyst regeneration at a maximum temperature of about 816 ^° C., but preferably not higher than 760 ^° C. The temperature in the second stage is the same or lower than in the first stage when the carbon on the catalyst is reduced to 0.05% or less, or even about 0.025% or less of the mass in the second zone. Indeed, the catalyst can be easily regenerated to such low carbon levels as 0.01% by this technique, even if the carbon content on the catalyst before regeneration was as high as about 1 wt. or more.

 Another particularly preferred technique for controlling or limiting the heat of regeneration transferred to the fresh feed from the recycled catalyst is to remove part of the heat carried by the recycled catalyst to additional material. The catalyst discharged from the regenerator is stripped with an appropriate stripping gas to remove oxygen-containing gases. Such steaming can, for example, be carried out at relatively high temperatures using a stream of nitrogen or an inert gas (s) as the stripping gas. The use of nitrogen or other inert gases is favorable from the point of view of eliminating the tendency to hydrothermal deactivation of the catalyst, which may occur as a result of the use of steam.

Feed (feedstock)
The present invention is applicable to the catalytic conversion of a feed of light gas oil or a feed of heavy residual oils, including bottoms of vacuum distillation and parts thereof, which have been subjected to preliminary partial hydrogenation to remove sulfur and nitrogen compounds from them and / or which have been partially decarbonized and demetallized by contacting with sorbent material under thermal visbreaking in the presence of a diluent and in the presence or absence of hydrogen. The sorbent material used in the visbreaking operation may be relatively inert or with such low catalytic activity that it is unsuitable for use in the catalytic cracking operation.

The process conditions used in the catalytic cracking operation of the present invention will vary depending on the composition and boiling range of the feed oil feed. Typically, the regenerated catalyst charged to the cracking operation in an upflow reactor should be at a temperature of 649-816 ^° C., typically 704-760 ^° C. The ratio of catalyst to oil and the partial pressure of the hydrocarbon feed will vary depending on the feed boiling range and volume used gaseous diluent, so that the products of the conversion of vaporous hydrocarbons containing a suspended cracking catalyst, an elevator gas and the supply of atomized dilution material s from the cracking zone of the reactor with an ascending layer at a temperature of 482-598 ^o C, preferably 510 566 ^o C.

 The invention consists in the following.

A gaseous refinery product, known as dry (lean) gas (usually obtained from a traditional gas concentration plant), consisting of at least 10% hydrogen and less than 10% C ₃ plus hydrocarbons, is used as the lift gas to form suspension hot freshly generated catalyst at a temperature of at least 704 ^o C. The dry gas is supplemented with steam and / or water as a heat rejecting heat agent in an amount sufficient to reduce the regenerated catalyst temperature to the desired low rate perature oil feed conversion. Preferably, the dry gas contains about 15-40 vol. more preferably 20-35% hydrogen and 2-8, more preferably O-6 vol. C ₃ plus.

A suspension of dry gas-vapor-catalyst is formed in the lower part of the reactor conversion zone with an ascending layer below the injection point of the oil supply and remains there for a residence time that prevents coke deposits on the catalyst in an amount of more than about 0.20 wt. or 0.25 wt. before contacting with a loaded oil feed boiling above about 343 ^° C. So, the residence time of a suspension formed from a dry gas-vapor-regenerated catalyst in an ascending bed reactor before contacting with a loaded oil is no more than a split second, but is sufficient to restore metal oxides on the catalyst to a less oxidized state or to a metallic state.

Refined product dry gas streams containing less than 10 vol. hydrocarbons With ₃ plus and 10-40 vol. hydrogen is recovered from aromatic desulfurization units under economically acceptable conditions that promote its use in the oil feed cracking operation of the present invention. Hydrogen sulfide in this dry gas stream is a valuable component for the formation of sulfide compounds with metallic impurities.

 Processing of carbometallic oil feeds on a catalyst containing up to 20,000 ppm Ni + V of metallic impurities accumulated on the catalyst can be carried out with improved product selectivity using the described dry gas in the case of one or more metal passivating agents.

Improved selectivity of the product is achieved by combining the processing of the present invention and contributing to a significant reduction in the paste of oil and coke, helping to reduce the high temperature of the catalyst regeneration operation in combination with regeneration steps exceeding the temperature of 760 ^° C. at each regeneration step.

In FIG. 1 shows a graph comparing hydrogen yields obtained using fatty recycled gas and dry recycled gas as an elevator gas for a catalyst before converting the oil feed into products boiling below 221 ^° C.

In FIG. 2 is a graph comparing the yields of C ₅ 221 ^o C gasoline obtained using fatty recycled gas and dry recycled gas as an elevator gas for a catalyst before converting the oil feed into products boiling below 221 ^o C.

In FIG. 3 is a graph comparing the selectivity to gasoline obtained using fatty recycled gas and dry recycled gas as an elevator gas for a catalyst before converting the oil feed into products boiling below 221 ^o C.

In FIG. 4 is a graph comparing the yields of coke obtained using fatty recycled gas and dry recycled gas as an elevator gas for the catalyst when converting the oil feed into products boiling below 221 ^o C.

In FIG. 5 is a graph comparing the reduction in the formation of C ₂ -minus by-products obtained using fatty recycled gas and dry recycled gas as an elevator gas for a catalyst before converting the oil feed into products boiling below 221 ^° C.

 In FIG. 6 is a hardware lift scheme comprising two regeneration stages near a cracking zone in an upstream reactor.

Below is the composition of the used "fat" gas (mol.). It should be noted that the C ₃ plus material and up to and including C ₅ hydrocarbons is about 45%
Hydrogen 19.74
Methane 17.72
Nitrogen 3.97
Carbon monoxide 0.72
Ethylene 5.93
Ethan 6.97
Oxygen + argon 0.03
Hydrogen sulfide Not determined
Propylene 12.48
Propane 4.16
Carbon dioxide 1.13
1,3-Butadiene Traces
1-Butene + Isobutylene 5.07
Trans-2-butene 2.59
Cis-2-butene 1.98
Isobutane 3.81
Normal butane 1.08
Carbon Disulfide Traces
Isopentane 2.27
Normal pentane 0.28
Recovered after n-pentane 10.05
Below is the composition of the used "lean" (dry) gas (mol.). It should be noted that the C ₃ plus material up to and including n-pentane amounts to about 8.68%
Hydrogen 20.01
Methane 33.10
Nitrogen 9.94
Carbon monoxide 4.03
Ethylene 12.07
Ethan 11.40
Oxygen + argon 0.13
Propylene 4.24
Propane 0.51
Carbon dioxide 0.01
1-Butene- + isobutylene 0.62
Trans-2-butene 0.17
Cis-2-butene 0.11
Isobutane 0.69
Normal butane 0.12
Isopentane 0.02
Extraction after n-pentane 2.19
In the experiments, residual oils containing carbometallic impurities were used, for which the Ramsbottom coke number, sulfur, nickel and vanadium content were determined, which are in contact with a typical liquefied cracking catalyst, which is a crystalline aluminosilicate with rare exchange metals (faujasite), followed by treatment during regeneration at elevated temperature, by a "dry" gas, a product enriched in hydrogen containing less than about 10 vol. C ₃ plus hydrocarbons and a fatty gas product enriched in hydrogen, catalytic cracking, consisting mainly of C ₄ and C ₅ hydrocarbons.

 The data obtained are presented in graphs.

In FIG. 1 shows that the upper curve representing the treatment of the catalyst with a fatty gas, corresponds to the production of a much larger amount of hydrogen during the subsequent conversion of the supply of residual oil to the catalyst suspension than was obtained using dry gas as an elevator gas to form a catalyst suspension at a temperature of 704 ^o C.

In FIG. Figure 4 shows that when using hydrogen-rich fatty recycle gas containing C ₄ and C ₅ hydrocarbons in significant quantities, significantly more coke is produced on the catalyst than was obtained when using hydrogen-rich dry recycled gas. The high deposition of coke contributes to high catalyst regeneration temperatures in excess of 760 ^o C.

 From FIG. 2 that the use of dry gas as an elevator gas provides a higher yield of gasoline product than was obtained when using fatty gas as an elevator gas. Therefore, the selectivity to the gasoline product is significantly improved.

 From FIG. 3 that the improved selectivity to gasoline corresponds to the curve obtained using dry gas for the initial contacting of the hot regenerated catalyst, compared with that obtained using hydrogen enriched fatty recycle gas.

 In FIG. 6 shows a reactor with an upward catalyst bed 1 with an expanding transition section in its upper part, which ends with a section with a large diameter 2 of the specified reactor. Particularly effective is the conversion of a loaded oil supply, such as the supply of residual oil, through inlet devices 3.4 or 5. The enlarged or large diameter section of the reactor with an upward layer 2 is provided with a plurality of nozzle devices 6 for supplying power near the upper edge of the transition section, which are used in a preferred embodiment, for loading oil feed. Vertically arranged inlet devices 3,4 and 5 provide the operator with significantly greater freedom in the time of contact of the dry gas-catalyst suspension inside the reactor with the ascending layer before separation of the resulting hydrocarbon vapor, catalyst and lift gas suspension.

With this hardware design, the hot regenerated catalyst at a temperature of 649-816 ^o C is first mixed with a dry elevator gas (liquefying gas) with the addition of steam and / or water as heat-removing material to form a suspension rising upward in a section with a limited diameter at a temperature suitable for carrying out catalytic cracking of a downward loading stream of residual oil fed through devices 4 or 5. So, feeding through inlet devices 3,4 and 5 with inlet dilution devices 6,7 and 8 allows you to significantly vary the atomization of the feed, the partial pressure and the contact time between the oil supply and suspended in a dry gas catalyst particles. In addition, the bottom section of the reactor with an ascending layer allows you to adjust the temperature of the regenerated catalyst by adding steam and / or water as a heat-removing agent along with lean skin gas,
The contact time between the supply of residual oil and the catalyst in the reactor with an upward catalyst bed depends on the composition of the feed and can be limited to the interval from 0.5 to 2 or 3 s when the oil feed is contacted with the catalyst at a temperature in the range of 704-760 ^o C to ensure a temperature of the exit from the reactor is 510-593 ^o C, preferably not higher than 566 ^o C. The reactor with an ascending catalyst bed can be of almost any vertical length, which will be compatible with the adjacent catalyst regeneration equipment, independently about whether the regeneration will be single or multi-stage, stripping of the catalyst and means for transferring the catalyst.

Steaming
The upper part of the reactor 2 with an ascending layer passes upward through the stripping zone 9 to form an annular stripping zone in the upper part of the large-diameter catalyst separation zone connected to the ring stripping zone. A stripping gas, such as steam or another suitable gas, is supplied to the lower part of the stripping zone through a conduit 10, flowing upward through it and counter-flowing to the catalyst particles moving downward.

Regeneration
The stripped catalyst is then sent via conduit 11 to catalyst regeneration, shown as a sequence of catalyst layers 12 and 13, regenerated in separate zones to remove carbon deposits of the conversion by burning without exceeding the temperature above 816 ^o C, preferably limited to 649-816 ^o C, usually the range of 704-760 ^o C.

Catalyst separation
A suspension of hydrocarbon vapor, an elevator gas catalyst and steam is withdrawn from the upper open end of the ascending bed reactor at a speed that should give a greater moment to the catalyst particles than the moments given to the vapor components, thereby establishing an upward trajectory at which the catalyst particles separate from vaporous material. A mixture of vaporous materials, often called gaseous material in the prior art, passes into an annular niche 14, into an open outlet channel in its upper part, and from there through a radial nozzle to a cyclone separator 15 at the outer end of each of these radial nozzles. Vapors separated from the entrained fine catalyst particles in cyclones 15 and 16 are fed through the connecting piping to the chamber 17 and the product is withdrawn through the pipe 18, directing it to fractionation and separation, which are not shown.

 Fine catalyst particles, separated in cyclones 15.16, are removed along the lower elbows, directing to the stripping and regeneration of the catalyst.

Catalyst
The hydrocarbon conversion operation under consideration, which should be carried out in the zone of the reactor with an ascending catalyst bed, is based on liquefied catalyst particles with a particle size of at least 10 • 10 ^-6 m (10 μm), usually the average particle size is 60-100 • 10 ^{- 6} m (60-100 μm), preferably below 85 • 10 ^-6 m (85 μm). Preferably, the catalyst is crystalline aluminosilicate or crystalline zeolite that has been exchanged for rare earths and / or ammonia to provide a catalytically active material that is dispersed in the matrix material. The latter may have catalytic activity. The catalyst, especially suitable for use in the method of the present invention, is a rare earth-exchanged crystalline faujasite-type zeolite having an open pore volume of the catalyst and an open pore size of the matrix, which will collect and / or accumulate significant amounts of metal impurities and significant cracking activity will remain and selectivity.

Supply (power)
An oil feed, such as a residual portion of crude oil, charged through inlet ports 3 and 4, may be mixed with steam and / or water, such as an acidic aqueous product, charged via pipelines 6 and 7. On the other hand, when loading an oil feed through the feed hole 5, the steam-water mixture can be introduced via line 8. The lower part of the reactor with the ascending layer 2 is supplied with dry elevator gas through the inlet line 19 to load the elevator gas with the formation of an upward moving suspension with hot regene particles catalyst, loaded into the lower part of the reactor through the pipeline 20. Dry elevator gas can be loaded into the reactor alone or in combination with steam and / or water introduced through the pipe 21.

The lower part of the reactor with an ascending layer of a limited diameter can be used to perform several different functions in addition to the formation of a moving upward suspension with the desired concentration of part of the catalyst in the range of 16-44 kg / m ³ . Namely, a passivating metal compound for passivation can be added to the elevator gas. Antimony can be added to passivate accumulated nickel deposits. Vanadium oxide can be passivated by combining hydrogen reduction to a lower oxide state, providing a high melting point of this oxide, alone or in combination with the addition of titanium, alumina and rare earth metals enriched in lanthanum. Therefore, no matter how the bottom of the bottom-up reactor is used prior to the injection of the atomized oil feed, it is essential for the ideas of the present invention that the use of a hydrogen-containing recyclable dry gas be consistent with a composition that is severely limited by the C ₃ plus dry gas components to a level that inhibits any significant deposition of coke on the catalyst, and before contacting with the supply of heavy oil to be cracked. The restriction of the hydrogen-containing dry gas to the C ₃ plus component is less than 10 vol. more preferably less than 8 vol. and even more preferably less than 6% improves the yield of gasoline, reduces the yield of hydrogen, increases the yield of light cyclic oil, and reduces the yield of butter and coke paste.

In a specific embodiment, dry cracked gas containing at least 15 vol. hydrogen, less than 10 vol. A C ₃ -plus of hydrocarbons in a mixture with water, taken in an amount sufficient to partially cool the regenerated catalyst to the desired low level of oil feed conversion before contacting with the atomized pre-heated residual oil loaded into the ascending layer of the dry catalyst gas suspension. The catalytic cracking of the loaded hydrocarbons in the fluidized bed is carried out at an upstream pressure above atmospheric pressure and the cracking operation in the upstream layer of the present invention can be carried out at a pressure of about 172 · 10 ³ - 1.137 · 10 ³ Pa. However, the partial pressure of the hydrocarbon atomized oil feed should be significantly lower due to mixing with the lift gas-steam and diluent atomized oil feed. Therefore, the partial pressure of the oil supply can be in the range of 27.6 172 • 10 ³ Pa, and the ratio of catalyst to oil can be in the range of about 5-15, more preferably 6-12, with careful contact between the catalyst particles and the atomized oil serve.

The upper part of the chamber of the regenerator has a larger diameter than the lower part of the chamber, and they are separated from each other by a distribution chamber of regenerating gas 22 located in the center and provided with an annular partition 23 directing the gas flow through channels 24. Many radial guiding means 25,26 from the chamber 22 is provided for introducing regenerating gas into the lower part of the catalytic layer 12 to be regenerated. Combustion-regenerating gas, such as air or oxygen-modified gas, is supplied via line 27 mixed with steam in the line at the desired oxygen concentration in an amount greater or less than that required to achieve partial removal of carbon deposits from the loaded catalyst particles, resulting in Combustion temperatures may be limited to the desired range. These gases are supplied through a chamber 22 and radial guide means 25, 26. In this first stage of catalyst regeneration, including the combustion of carbon deposits, carried out in the presence of oxygen, carbon dioxide and steam, the regeneration temperature is preferably kept low in the range of 593-871 ^o C preferably 649-815 ^o C, more preferably in the range of 690-760 ^o C.

 Partially removable carbonaceous material is removed from the catalyst bed 12 under conditions of production of flue gases enriched in CO and containing carbon dioxide, sulfur, nitrogen and water vapor. Thus, the generated flue gases pass through one or a combination of several cyclones represented by cyclones 28 and 29 to remove entrained fine catalyst particles recovered through the provided downstream elbows. The flue gases are then sent from the cyclones 28 and 29 to the chamber 30 for their removal through the pipe 31. Such CO enriched flue gases usually enter a CO boiler (not shown) for the recovery of process steam.

The partially regenerated catalyst contained in the layer 12 is removed from the lower part as a downward flow through the remote cooling zone of the catalyst 32 during indirect heat exchange in the tubular heat exchanger 33. High pressure steam of the order of 3.1 • 10 ⁶ Pa is generated and discharged through the pipe 34, boiler water is charged via line 35 into the distribution chamber at the bottom of the cooler 36 associated with the heat exchanger tubes 33. The partially cooled in chamber 36 when the catalyst temperature is typically 28-111 ^o C 55-83 ^o C, is withdrawn and sent via Pipelines row 37 in the catalyst layer 13 located on the second stage catalyst regeneration in chamber 7. The riser 38 connects the layers 12 and 13, is provided to direct catalyst directions without cooling from the upper layer to the lower layer (if necessary). However, the main or primary catalyst stream between the layers passes through a cooler 32 to maintain the desired catalyst temperature maintained in a sequential regeneration system. The temperature maintained in the second stage, which is a layer 13, is limited to a range of 649-816 ^o C, more preferably 704-760 ^o C. The temperature of the regenerated catalyst in a dense liquefied layer 13 may be equal to, higher or lower than the temperature maintained in a dense liquefied layer catalyst 12 in the first stage of regeneration of the catalyst. In one embodiment, the amount of air or oxygen-modified gas charged into the catalyst bed 13 through line 39 and passing through the grate 36 may be equal to or greater than the amount required to completely burn the remaining carbon on the partially regenerated catalyst. In this case, flue gas enriched in CO _{2 is formed} , which may contain a certain amount of unspent oxygen. Preferably, the flue gas leaving the upper dense phase of the catalyst layer 13 is free of combustion-supporting amounts of CO to prevent possible subsequent combustion. The CO ₂ enriched flue gas from the second stage of catalyst regeneration at elevated temperature passes through openings 24 in the baffle 23 to the lower part of the layer 12 for mixing with the regenerating gas charged through the distributors 25,26, thus transferring heat to the first stage of catalyst regeneration . All flue gases that are combustion products in the second stage of catalyst regeneration to reduce the coke residue to 0.05 wt., Or lower, for example, to 0.01 wt. coke on the particles of the regenerated catalyst is passed through the catalyst bed 12 of the first stage of regeneration. The regenerated catalyst is diverted to the bottom of the catalyst bed 13 through a pipeline to the bottom of the reactor with an upward catalyst bed 1.

 The table shows the data of the process in various operating conditions and the results obtained.

 Thus, the data of the table and the graphic illustrations show that the method according to the invention improves the selectivity with respect to the formation of gasoline and light cyclic oil.

Claims (1)

A method of processing a petroleum oil fraction by catalytic cracking in the presence of a cracking catalyst containing crystalline zeolite in the cracking zone of an upstream catalyst bed using a hydrocarbon lift gas to form an ascending layer of a suspension of catalyst particles and then contacting the suspension with the feed at a hydrocarbon conversion temperature to obtain target products, characterized in that, in order to increase the selectivity of the process with respect to education target gasoline and light cyclic oil, gas of oil refining processes containing at least 10 vol. hydrogen and less than 10 vol. C 3 ₊ hydrocarbons with a contact time between the lift gas and suspended catalyst particles of 0.01 0.5 s and maintaining the coke content on the catalyst with the lift gas of not more than 0.20 wt.

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