CS227651B2 - Method of cyclic fluid catalytic cracking - Google Patents

Abstract

1499682 Fluid catalytic cracking STANDARD OIL CO 4 March 1975 [4 March 1974] 8885/75 Heading C5E [Also in Division B1] In a cyclic catalytic hydrocarbon cracking process wherein deactivated catalyst is withdrawn from the reactor, stripped, regenerated and recycled to the reactor, the regeneration comprises (a) fluidizing the deactivated catalyst, in physical association with a carbon monoxide-oxidationpromoting catalyst of a metal of At No. of at least 20, with a molecular oxygen-containing gas so as to burn substantially all coke from the catalyst and convert substantially all of the carbon monoxide produced to carbon dioxide, most of the heat so produced being absorbed by the catalyst particles and the temperature being maintained at from 1000-1500‹ F.; and (b) withdrawing an effluent gas containing not more than 1À0 vol. per cent of carbon monoxide and regenerated cracking catalyst containing absorbed heat and having no more than 0À10 wt. per cent residual coke. The initial coke combustion and the carbon monoxide oxidation may be effected in a single dense phase bed or the former may be effected in a dense phase bed and the latter in a dilute phase bed. The oxidation catalyst may be metallic or oxidic and may contain Zn, Ce, U, Ti, Zr, Bi, Cr, Mo, W, Mn, Fe, platinum metals or heteropoly oxides. It may be present, either supported or unsupported, in physical admixture with the cracking catalyst or may be impregnated thereon. The cracking catalyst may be a crystalline aluminosilicate in a silicaalumina matrix. The cracking process may use a feed boiling above the gasoline boiling range, e.g. at 400-1200‹ F. such as light or heavy gas oils, wide-cut gas oils, vacuum gas oils, kerosenes, decanted oils, residual fractions, reduced crude oils or cycle oils derived from these or fractions from shale oil, tar sands processing, synthetic oils and coal hydrogenation. Suitably, cracking is effected exclusively in a 'riser' reactor at 950-1000‹ F. with a contact time of 3-10 secs., a throughput ratio of 1 to 2, and a catalyst to oil weight ratio of 2 to 10. Spent catalyst may be stripped at 850-1050‹ F. with steam, nitrogen or flue gas at 10-35 p.s.i.g.

Description

BACKGROUND OF THE INVENTION The present invention relates to a cyclic fluid catalytic cracking process wherein hydrocarbons are cracked in a fluidized state by contact with catalyst particles in the reaction zone, the cracked catalyst deactivated catalyst is separated from the hydrocarbon cracking reaction zone and devolatilized the volatile-free catalyst is passed to a regeneration zone and fluidized therein with a regeneration gas containing a molten crucible to form a lower dense catalyst particle phase and an upper dilute catalyst particle phase, the coke is burnt from the catalyst particles by contact with the regeneration gas wherein carbon monoxide and carbon dioxide are formed as combustion products, the gaseous product and the regenerated catalyst are separately taken from the regeneration zone and the regenerated catalyst is recirculated to reaction zone for further use in hydrocarbon cracking.

* Catalytic cracking of heavy mineral oil fractions is one of the predominant refining operations used to convert crude oil to the desired fuel products, such as high-octane gasoline, used in internal combustion and spark ignition engines. Examples of fluid catalytic conversion processes are fluidized bed cracking processes where high molecular weight liquid or gaseous hydrocarbons are contacted with hot finely divided solid particles. of the catalysts, either in a fluidized bed reactor or a longitudinal fluidized bed reactor, and the catalyst / hydrocarbon mixture is maintained at elevated temperature in a fluidized or dispersed state for a time sufficient to achieve the desired degree of cracking to lower molecular weight hydrocarbons typically used in engine fuels and destiolated fuels. Suitable hydrocarbon feedstocks are those which generally find a boiling point above the boiling point of gasoline, for example from about 203 to about 650 ° C, and typically crack at temperatures of about 470 ° C or 570 ° C.

In catalytic processes, a certain amount of non-volatile carbonaceous material or coke is deposited on the catalyst particles. The coke contains highly condensed aromatic hydrocarbons which contain a minor proportion of hydrogen of about 4 to 10 hmoo. %. When a coke coating is formed on the catalyst, the cracking activity of the catalyst is reduced and is only selectively for producing gasoline or hydrocarbons capable of being blended with β gasoline to drive the moor. The catalyst particles can recover most of their original abilities by removing most of the coke adhering to them by an adequate recovery process.

Catalyst regeneration is accomplished by burning coke coatings on the catalyst surface using a mooelkillary oxygen-containing gas as air. Many regeneration processes are carried out on an industrial scale, whereby a significant recovery of catalyst activity is achieved depending on the degree of coke removal. As the coke is gradually removed from the catalyst, the remainder of the coke becomes more difficult to degrade and is virtually accepted as an economic compromise when the level of recovered catalytic activity is reduced.

The combustion of coke coatings from catalysts requires a large volume of Donkey or air. Oxidation of coke in a simplified way can be known as carbon oxidation and is represented by the following of the chemical equation:

(and) C + o ₂ ------- ------- ^C0 2 (b) 2C + ° ₂ - 2C0 (C) 2CO + ₂ ------ ------- 2C02

Both reactions under (a) and (b) occur under typical catalyst regeneration conditions, where the catalyst temperature can range from about 630 ° C to about. 710 ° C and is a typical chemical action between gas and solid, whereby the catalyst regenerates at temperatures within this range. The effect of each temperature increase is reflected in the increased rate of carbon combustion - and more completely - the removal of carbon - or coke from the catalyst particles. When peace rises. combustion, higher heat dissipation increases if it is sufficiently sufficient to multiply! and then the gas phase reaction of (c) occurs. This reaction is induced and supported by free radicals.

Often, a difficult situation, which is undesirable in practice, often occurs when regenerating the fluidized bed catalysts, a phenomenon known as post-combustion is described, for example, in Hengstebeck, Petrov Processing, McGraw-HHll Book Co., 1959, p. 175 to 175 and which is further discussed in Oi! and Gas Journal, - Vol. 53 (No. 3), 1955, pp. 93-94. The term refers to the further combustion of carbon monoxide to carbon dioxide as represented by the above reaction (c) and which is highly exothermic.

This further combustion has to be significantly eliminated in catalyst regeneration processes, since this could lead to very high temperatures, which could damage the apparatus and cause permanent deterioration of the catalyst particles. Many processes for the recovery of fluidized bed catalysts have been obtained by testing and further combustion, and a number of significant processes have been developed regarding a number of means for controlling the regeneration processes to prevent this post-combustion. Later - for various reasons, efforts have been made to increase the regeneration temperatures - for various reasons. A number of processes have also been developed to control the temperature in the regenerator at the time of post-combustion start using a variety of control measures and to control the feed to the recovery vessel, such as, for example, U.S. Patent Nos. 3,161,583, 3,166,393 and 3,513 08 * 7. In typical practice today, therefore, they usually achieve very little oxygen and substantial potentials of carbon monoxide and carbon dioxide in near equimolar amounts to prevent the subsequent appellation of the flue gas from the catalyst regenerator-catalyst.

Further combustion of carbon monoxide to carbon dioxide is an interesting source of thermal energy because the reaction under (c) is highly exothermic. Further combustion can take place at a temperature of about 595 ° C and releases approximately 1100 kcal per 0.45 kg of carbon monoxide which has been oxidized. This typically represents about 1/4 of the total heat that can be generated by coke combustion.

The combustion of the carbon monoxide can be carried out in a controlled manner in a separate zone or by using a boiler heated with carbon monoxide to release it. the release of gas from the catalyst, such as described in U.S. Patent 2,753,925, wherein thermal energy is used in various refining operations, such as in the production of high pressure steam. Other uses of this thermal energy are described in U.S. Pat. Nos. 3,012,962 and 3,137,133 (turbine propulsion) and U.S. Pat. No. 3,363,993 (preheating oil). In these heat recovery processes, a separate and complex device is needed, but this does not contribute to minimizing the release of carbon monoxide into the atmosphere as a slogan of combustion gases, and in this process.

it does not prevent the potentially serious risks of contamination by the atmoofers.

Silica and alumina catalysts, commonly used for many years for cracking petroleum hydrocarbons, are not particularly sensitive to the residual coke rate on the catalyst, provided that the amount of coke is not more than about 0.5 wt. %. However, the capacitors have been largely replaced by catalysts in which crystalline aluminosilicate, known as zeolites or grit sieves, are additionally incorporated.

The catalysts containing the moiocecular sieve are much more sensitive to the residual coke rate and in the presence of 1 coke are severely adversely affected both in terms of catalyst activity and catalyst selectivity for converting the feedstock into. the desired product or products. Due to the difficulties in the regeneration processes of conventional catalysts in removing the last increase in residual carbon in practice, the level of coke coating corresponds to the residual carbon. coke on regenerated. from about 0.2 to about 0.3 wt. %.

Due to the increased activity and selectivity achievable with low-level coke coating coke catalysts. is a very interesting factor in the discovery of means to further reduce residual coke coating levels. It is highly desirable for the coke coating to be below about 0.05% by weight, but usually this cannot be achieved by means useful in industrial practice. Considerations relating to larger recovery vessels, larger catalyst movement devices, greater heat loss, and the like are all facts that prevent such ideal and most of the catalyst activities from being achieved.

There is also known a process which relates to an improved cracking process including an improved process for the regeneration of catalysts used in the conversion with fluidized bed catalysts, hydrocarbons, wherein the catalyst is deactivated by coke deposition on the catalyst surface. The process allows the level of coke coating on the recovered catalyst to be smoked at an extremely low level, while at the same time favoring a favorable thermal equilibrium in the conversion unit, wherein the combustion gases have an extremely low carbon monoxide content. According to one embodiment of the method of the invention described in said patent, the combustion of carbon monoxide to carbon dioxide is carried out substantially to completion in a recovery vessel in a relatively diluted secondary catalyst regeneration zone, preferably at a temperature between approximately 650 and 81 ° C. . between approximately 680 and 780 ° C. The temperature in the secondary zone is usually at least about 10 and 38 ° C higher than in the first regeneration zone. The partially regenerated catalyst from the relatively dense primary zone for catalyst regeneration can be directed to the secondary zone in a plurality and at high speed in a controlled manner, if sufficient to absorb all the heat released by the combustion in the second (i.e., secondary) zone. Although most of the coke adhering to the catalyst is combusted in the primary zone, additional coke from the partially regenerated catalyst in the secondary zone and the substantially coke-free catalyst can be recovered for recirculation to the hydrocarbon conversion zone.

Heat generated by the combustion of carbon monoxide and absorbed. the regenerated catalyst provides part of the heat required in the hydrocarbon conversion zone. In addition, the flue gas stream released from the secondary recovery zone is in the undersize free of carbon monoxide. According to another embodiment of the method described in said patent, the entire combustion, including oxidation of coke or carbon on the catalyst and oxidation of carbon monoxide, takes place in one relatively dense regeneration zone depending on the actual control of the initial regeneration temperature and gas velocity.

The problems of the prior art are eliminated by the cyclic-fluidized catalytic cracking process, which is based on the fact that coke and carbon monoxide are burned in a regeneration zone up to a content of not more than 0.10 wt. % of the residual coke in the regeneration catalyst and containing not more than 10% by volume of carbon monoxide in the gaseous product, the regeneration zone being maintained at the regeneration zone. at a temperature of from 537.77 to 815.55 ^V C, the regeneration gas and is fed to the regeneration zone in a molar-flow potential. of oxygen in excess of about 0.1 to 25% of the stoichiometric amount required for complete conversion of coke to carbon dioxide and steam, in the regeneration zone, a carbon monoxide-promoting agent containing one or more metals of at least 20 and selected from from groups IB, II-B and III to VIII of the Periodic Table, a portion of the heat generated by the combustion of carbon monoxide is absorbed by the catalyst particles and the absorbed body in the catalyst particles is maintained in the catalytic cracking process by re-feeding the heated regenerated catalyst particles to the catalyst. .reaction zone.

According to the invention, a regenerated catalyst for the conversion of very low coke hydrocarbons is obtained, while monitoring that the amount is less than about 0.05 wt. % and preferably in the rd range of about 0.01 to 0.03 µm. %, which is achieved by controlling substantially complete combustion of the rat. together with the substantially complete combustion of the carbon monoxide gas present or formed during catalyst regeneration.

• The method is characterized. in that the oxidation catalyst is in physical association with the hydrocarbon conversion catalyst and the oxidation catalyst may be in the form of a solid metal oxide containing material. In the process of the invention, regeneration temperatures are used which are sufficiently high to activate the catalyst containing the solid metal oxide to induce oxidation of the carbon monoxide and to achieve complete combustion of the deposited coke layers. on conversion catalysts.

The temperatures are not so high that the conversion catalyst particles are not deactivated by thermal action or that the safety of the recovery vessel or parts thereof is compromised, or that the recovery vessel does not become inoperable. The regeneration temperatures may preferably be from about 380 ° C to about 660 ° C, preferably from about 400 ° C to about 600 ° C. These combustion temperatures are achieved with the aid of the oxidation catalyst as described below.

A major advantage is that a regenerated catalyst is obtained which generally has increased levels of activity and selectivity very close to those of fresh conversion catalysts, in particular. Conversions, which are carried out in very short time limits in ascending reactors. In this manner. higher conversions of the feedstocks and higher yields of the monitoring and conversion can be achieved. The oxidation catalyst assists in stabilizing controlled regeneration by treating the carbon monoxide. Furthermore, the oxidation of carbon monoxide is stopped by decreasing the temperature or by increasing the gas flow rate, which can lead to expulsion or reproduction of the carbon monoxide content. It is substantially removed because the oxidation of carbon monoxide can be carried out catalytically.

In regeneration processes using a lower dense phase zone and an upper dilute phase zone in the regenerator, oxidation. The carbon monoxide to carbon dioxide can be carried out to a greater extent, often at least about 60%. and often about 65 to 95%, or more, so as to be in the dense catalyst phase in the fluidized bed catalyst regenerator.

The dense phase oxidation of carbon monoxide to carbon dioxide provides heat to aid in the combustion of coke coatings on the catalyst catalyst. When a substantial portion of the carbon monoxide is oxidized in the dense phase, the carbon monoxide molarity is present for the upper phase combustion of the fluidized bed catalyst in the regenerator, thereby substantially reducing or avoiding further combustion and higher temperatures due to uncontrolled over-combustion of carbon monoxide at the top of the fluidized bed catalyst in the regenerator, which could be detrimental to the material used in the co-construction of the waste gas reactor and to the particulate matter collectors in the waste gas, such as cyclones to generate heat also damage the activity of the catalyst.

Carbon monoxide content of waste gases. The resulting regeneration process can be maintained at a level of less than about 0.2 vol%, for example in an amount of about 500 to 1000 vol, ppm. Preferably, the content is kept at a lower level, for example in the range of from 0 to about 500 parts per million. This low concentration of oxide. carbon monoxide in the flue gas stream allows them. 'Directly discharged into the air while adhering to standards concerning air purity' 'This advantage. obtained by the invention further allows to eliminate the bases that are otherwise required to provide various additional carbon monoxide incineration equipment and associated different turbine devices or other partial energy recovery devices produced by the afterburning of carbon monoxide.

The novel regeneration method according to the invention is preferably used as a precaution in cracking catalytic converter utilizing. Fluidized bed catalyst, especially where a significant part of the conversion is carried out in a reactor system using dilute transfer phases where the catalysts need to be very active at relatively high space velocities. Low coke coating - particularly on the recovered catalyst. preferably when using fluidized-bed cracking catalysts containing catalytically active crystalline aluminum silicas, otherwise referred to as. zeooity or mooelouLear sieves ·

The cracking activity of the catalysts containing the molar sieve and their selectivity for converting the hydrocarbon feedstocks to the products of interest is very strongly influenced by the increased deposition of residual carbon or coke on the catalyst during regeneration. A particularly preferred hydrocarbon conversion process for the practice of the present invention includes a catalytic fluidized-bed catalyst cracking process to convert gas oils and heavier hydrocarbon feedstocks to lower-temperature hydrocarbon components suitable for blending into combustion engine, jet engine, for dommies and industrial heating and the like.

It has also been found that processes using a cracking catalyst, particularly a molecular sieve type catalyst, are very advantageous when an oxidation promoting catalyst is additionally used to obtain a regenerated catalyst with a very low coke content producing useful heat. The level of coke coating on the catalyst particles is less than about 0.05 wt%, and preferably about 0.01 to 0.03 wt%. The process relates to deserts of a cracking catalyst in association primarily physical rather than chemical with a catalyst promoting oxidation of a system that promotes substantially complete combustion of carbon monoxide. From this process, it is possible to obtain combustion gases between carbon monoxide, less than about 0.2 vol%. %, for example about 500 to 1000 ppm and even as low proportions as about 0 to 500 ppm. The process according to the invention also relates to the provision of how the heat recovered is recovered directly to the conversion catalyst. especially in the recovery vessel. All devices used in this procedure are usable. in carrying out the process of the invention and a device directed to substantially complete combustion within one relatively dense phase, it is particularly preferred.

Suitable cracking catalysts in this process include (silicon dioxide and / or alumina catalysts), including acid type catalysts, and catalysts may contain other refractory metal oxides such as magnesium and zirconium. Preferred cracking catalysts are those containing crystalline aluminosilicates, Crystalline aluminosilicates typically have a silica silica to alumina content of at least about 2: 1, for example about 2 to 12: 1, preferably about 4 to 6: 1. the aluminosilicates are usually obtained or produced in sodium form, and this component can be reduced, for example, to less than about 1 and about 4 or even less than about 1 wt% by ion exchange with hydrogen ions or ions having the hydrogen polymers such as ammonium or polyvalent metal ions Suitable polyvalent metals are calcium, strontium, barium, rare earth metals such as tare, lanthanum, neodyminium and naturally occurring rare earths and mixtures thereof ·

Useful crystalline masses are those capable of maintaining their porous structure under high temperature conditions during catalyst production, hydrocarbon processing, and catalyst regeneration. Crystalline aluminosilicates often have a uniform pore structure of extremely small size. The pore cross-section is in the range of about 6 to 20%, preferably about 10 to 15%. Silica based catalysts having a larger proportion of carbon dioxide, for example those containing about 60 to 90 wt. % of silica and about 10 to 40% by weight of alumina are the preferred base for the production of catalysts.

the catalyst particles are finely divided, for example, they may have an average particle size in the range of about 20 до or less than about 150 до, i.e. dimensions that are in a form suitable for fluidization. For example, the fluidizing gas in the dense zone of the regeneration device may have a velocity in the range of about p, 06 to 1.2 m per second, preferably 0.15 to 0.9 m per second.

The dense bed regeneration gas contains free or molecular oxygen and the oxygen is preferably transferred to the recovery plant in an amount slightly higher than that required to completely burn the coke, i.e. carbon and hydrogen, to carbon dioxide and steam. the excess over the amount required to completely burn the coke may vary between about 0.1 to about 20 or more% of the theoretically required stoichiometric amount of oxygen to completely burn the coke, but preferably not more than about 10%. A 10% excess of air gives only about 2 vol% oxygen in the outgoing flue gas stream.

Preferably, the concentration of molecular or free oxygen and carbon monoxide at any point in the regenerator is maintained outside the explosive range under these conditions, preferably the concentration of carbon monoxide is below the explosive limit under these conditions to avoid any risk of explosion.

The regeneration gas, in addition to the free or molecular oxygen, may contain an inert or diluent gas such as nitrogen, steam and the like, circulating gas from the regenerator effluent, etc. Often, the oxygen concentration in the regeneration gas at the inlet to the regenerator is approximately preferably about 5 to 25 vol%. Since air is preferably used as the oxygen source, a larger proportion of the inert gas is nitrogen.

The inert gas can serve to dissipate excess heat from the coke combustion to the catalyst. The hot inert gas source is the off-gas from the recovery plant, and part of the gas may be recycled to the recovery plant and, for example, combined with a sufficient amount of incoming air or other oxygen-containing gas that necessarily contains pure oxygen to produce the necessary oxygen content. In this way, the circulating gas can be used with direct heat exchange to raise the temperature of the regeneration gas and thus achieve further heat savings in the system.

According to the invention, the hydrocarbon cracking catalyst and the solid oxide-promoting catalyst 1 are in the recovery vessel. The solid oxidation catalyst may be impregnated into the hydrocarbon cracking catalyst structure or may be impregnated into the hydrocarbon cracking catalyst structure or may be introduced into the catalyst base mass during its production. As an alternative ae, the solid oxidation catalyst can be produced in finely divided form as a powder, separated from the hydrocarbon cracking catalyst or can be supported on another substrate and mixed with the hydrocarbon cracking catalyst. In all these processes, it is necessary to avoid ion exchange of the basic component of the oxidation catalyst with the zeolite.

The oxidation catalyst carrier, if present, may be at a lower degree of catalytic activity or even inert to the oxidation and hydrocarbon conversion reaction. It may, for example, be ceramic in nature. Desirably, the carrier is porous and often has a specific surface, including a pore area, of at least about 10, preferably at least about 50 m‘ / g. Examples of carriers of this type are silica, alumina, combinations of silica and alumina, and the like.

»

The solid oxidation catalysts used according to the invention may be of the type used and commonly known to those skilled in the art to increase the oxidation of carbon monoxide in the presence of molecular oxygen. These catalysts generally contain a catalytic metal which favorably influences oxidation, and the metal may be in combined form as an oxide or may be of an elemental nature.

Common catalytically active metals are heavy metals according to the Periodic Table of the Elements, and in groups Ib, LLi, - or III to VIII and having an atomic number above about 20. Generally the most active oxidation catalysts are platinum group metals such as platinum, palladium and rhodium or metals from groups IB, IIB, V, VI as molybdenum and tungsten, from - group - VII, iron, further from group VIII as copper, chromium, nickel, - msangan, cobbat, vanadium and the like, rare earth metals like cerium and ytterbium, uranium and the like. Preferred solid oxidation catalysts may comprise two or more catalytically active metals, combined physically or chemically. For example, the metal chemistry is a bimetallic or polymetallic salt such as an oxide, for example derived from tungsten and cerium, tungsten and boron, tungsten and aluminum, tungsten and iron, tungsten and germsn, molybdenum and iron, molybdenum. and cerium, molybdenum and germanium and the like. Examples of catalytically active metal combinations that can promote the oxidation of carbon monoxide without adversely affecting the hydrocarbon cracking process are iron and manganese oxides, iron and rhenium, iron and cerium and the like.

One method of preparing the solid oxidation catalysts of the present invention is to impregnate a suitable water or organic solvent soluble salt carrier derived from a catalytically active metal. The impregnation can be carried out in any manner that does not destroy the structure of the carrier. Preferably, water-soluble nitrate salts are used in the impregnation solution, since the remainder of the thermal decomposition of the nitrate salts is relatively harmless to the activity of the hydrocarbon cracking catalysts.

Halogen salts or metal sulphates may be used for impregnation, however, since during the thermal decomposition of the salt halogen or sulphide may be formed, which may be detrimental to the hydrocarbon cracking catalyst activity, these impregnation processes are used when the oxidation catalyst is deposited. to a separate substrate that does not significantly adversely affect the hydrocarbon cracking reaction. The catalyst may be stashed and inerunir supported. against, a cracking reaction, or can be impregnated to a lesser extent to cracking. hydrocarbon.

Np '. · ^^ · ^}! differs, for example, from the cation exchange in that the metal is. in the anionic portion of a compound soluble in water or in an organic solvent. The impregnation results in greater deposition on the surface of the catalyst, in contrast to the ion exchange there is greater diffusion and therefore less deposition on the surface. During impregnation the metal is deposited on the surface, there is no significant ion exchange between the metal and the carrier.

β

When impregnating the carrier, the metal which promotes the oxidation of carbon monoxide may be present as a water-soluble salt or in an organic solvent solution in an amount sufficient to retain the desired amount of metal on the carrier and contact the carrier with the solution. reduce к (Removal of solvent while leaving metal deposited on the surface.

Further heating may be required to convert the metal to the active state, such as calcination or heating in the presence of hydrogen, or another reaction atmosphere, or an inert atmosphere. Alternatively, part or all of the desired activation can be carried out in the recovery vessel during regeneration. · In general, to achieve the best distribution of the metal compound on the catalyst, it is recommended that the metal compound solution be as concentrated as possible if practicable.

Another method for depositing a catalytically active metal on a surface, especially a porous surface, such as crystalline aluminosilicates, is by adsorption of a liquid degradable metal compound on the surface, followed by thermal or chemical decomposition of the metal compound. The substrate may be activated by heating to remove adsorbed water and then contacted with a liquid degradable metal compound to effect adsorption of the compound into the substrate. Typical such compounds are metal carbonyl compounds, alkyl metal, volatile metal halides and the like. The adsorbed compound can then be reduced thermally or chemically to the elemental metal, leaving a uniform dispersion of the active metal promoting the oxidation of carbon monoxide to carbon dioxide on the substrate surface. The thermal reduction can then be carried out in the recovery vessel during the recovery process.

Another method of supplying a catalytically active metal that promotes the oxidation of carbon monoxide to a support comprises admixing the active metal component to a substrate forming material, for example, silica gel or silica-aluminum gel prior to spray drying, or other physical forming process, and before drying the mass for the preparation of the oxidation catalyst · The resulting catalyst body can then be calcined to form a catalytically active material · As a variation, the heat treatment can be carried out in a catalyst regenerator in a cracking system.

The substrate for the solid oxidation catalyst may be a hydrocarbon cracking catalyst, or a portion thereof, or may be different therefrom, for example a non-catalytic, porous, solid support. If the hydrocarbon cracking catalyst has the function of a substrate, care must be taken in selecting the process that the cracking activity and selectivity of the catalyst are not adversely affected. Preferably, if the hydrocarbon cracking catalyst is of the type having ion exchange sites, the ion exchange is completed before the oxidation catalyst is deposited.

The amount of oxidation promoting metal to be used to promote the oxidation of carbon monoxide should be as small as possible to support the oxidation of interest, often about 0.01 to 5% by weight. %, preferably 0.05 to 1 wt. % based on the amount of hydrocarbon cracking catalyst used.

It is preferable that instead of depositing the oxidation catalyst on a support, a solid oxidized shell catalyst is used as a catalytically active metal oxide which is admixed with a hydrocarbon cracking catalyst. In general, the advantage of a powdered metal oxide oxidation catalyst is that the catalyst is relatively inexpensive that is readily introduced and handled in the fluid catalyst system in the reactor and regenerator.

The particle size of the oxidation powder catalyst should be selected so as to avoid particle separation at fluid velocity. Desirably, the metal oxide particles were not so small as to cause the difficulty of over-carrying the catalyst with the gas phase from the bed as entrained particulate matter.

however, the filto, cyclone precipitation device · likewise together with the β workflows used in fluidized bed catalysts to recover as much entrained particulate matter as possible and return it to the system to reduce losses. The powdered metal oxide mixes to provide sufficient strength to avoid excessive abrasion and dusting of the powder particles. Typically, the particle diameter of the metal oxide catalyst is from about 0.5 or from 1 to 100 μα, preferably less than about 50 µm. It should be noted that particles of small sizes, i.e. particles with a diameter of less than about 1 µm, for example about 0.01 to 0.5 µm, may tend to form clusters of larger size, which may preferably be used in the process of the invention. As an example, powdered kyeiičníků the metals which can pouužt to promote CO oxidation, the iron oxide, ferrous further ^ sličník 'ferric together with zinc oxide, as kreliční ^to manganese ⁿ $ ^dioxide, křsHčn'', cerium, ^p ODO ^b No. ;

1 and 2 show, in a vertical view, partly in section, an embodiment of a device suitable for regeneration according to the invention. Regeneration of the catalyst used, in particular of the molecular sieve type, in various hydrocarbon cracking processes where the catalyst is physically associated with an oxidation catalyst, as mentioned above, can be carried out. to carry out the process according to the invention. This improved process can advantageously be used in a number of already existing petroleum hydrocarbon cracking plants, particularly in plants employing different spatial arrangements, separation and regeneration in the cracking of fluid catalysts β.

Giant. 1 illustrates one embodiment of a recovery apparatus according to the invention, wherein a lower inlet of separated inactive catalysts coming from a cracking reactor (not shown) to a recovery apparatus is used. Only the catalyst from the separation zone connected to the catalyst outlet from the reactor enters the bottom of the regeneration vessel ²⁸ in physical association with the oxidation promoting catalyst.

The catalyst flows upward through the inlet ducts 2 and j which flow into the dense catalyst bed through the heads 4 and 2. The dense phase catalyst bed is held in the lower section 4 of the recovery vessel 2 and reaches up to the edge of the catalyst phase 2. The dense phase bed canalizer is fluidized by the combustion air flow through line 8 through valve 2 and via line 10 to the air ring 11.

A substantially balanced air flow through the regeneration zone could be achieved if additional air rings (not shown) were used as needed. The combustion of the coke contained on the catalyst used in the presence of air begins in the bed forming a dense phase. Higher temperatures can be achieved by temporarily burning the injected heavy oil fed as a fuel to the burner nozzle and combusted inside the recovery vessel 2. For this purpose, cracking residues after the de-heat can be used, withdrawn as a bottom stream from the fractionation tower. Said heavy oil may be fed via line 12 through vents 13 and 14 which terminate in the nozzle, placed in the air ring 11. The fluidizing air velocities continuously carry part of the catalyst particles up to the dilute phase zone which fills the upper part 15 of the recovery vessel 2, that is, the section above the limit of 2 phases of the catalyst. Coke combustion continues in the dilute phase zone and the flue gas together with the entrainment catalysts are withdrawn into the cyclone training separators 20 and 21 forming the first stage of the cyclone separators. - ^z

Most of the catalyst particles are separated in first stage cyclones and removed downwardly through tubes 22 and 23 immersed in the dense phase zone. The gases and remaining catalyst particles are passed through the conduits 24 and 25 through the cyclone inlet to the second stage cyclone separators 26 and 27, where substantially all of the remaining catalyst is separated and discharged down through lines 27 and 29 immersed in the dense bed. .

substantially all of the gaseous combustion product is then discharged via lines 30 and 31 to vessel 11 and finally emptied from the recovery vessel 2 via line 33. This gaseous effluent can be used in a suitable manner to exchange heat (not shown) through the refinery feed stream or the like. . The regenerated dense bed catalyst is discharged via lines 34 and 35 equipped with header 36 and 37 and returned to the cracking reactor.

Although the combustion air supply normally constitutes an excess of oxygen over the amount required to achieve complete combustion of the coke on the catalyst particles to produce steam and carbon dioxide, the coke in the dense bed must not always be completed in the present embodiment. In this case, the combustion gases exiting from the bed of dense phase thus contain substantial carbon monoxide as well as carbon dioxide and nitrogen. .

The remaining coke on the catalyst and the carbon monoxide - is essentially completely burned in the dilute phase zone, generating a large amount of heat. When the carbon monoxide is combusted in the dilute phase, a high temperature band usually forms in the - most of the dilute phase band - and especially around the X-marked site and can be easily observed by the window shown in that plane.

The recovery temperature control within the diluted phase is accomplished in part by the absorption of heat - into the mass of catalyst particles that are carried upwardly by the flow of combustion gas or carried upwardly from the solid phase bed through the discharge tube 40 and distribution head 41. the catalyst into the dilute phase zone. The catalyst can be discharged by air, steam or other inert gas entering line 42 through the outside. 43 and a nozzle 44 that reaches the lower end of the discharge tube 40 at a short distance therefrom. The excessive temperature level at the top of the recovery vessel 1 can also be controlled by steam supply, for example via lines 45 and 46 via outlets 47 and lines 48 to the distribution device 49. steam. The temperatures near the vessel 32 can also be controlled by the gentle inlet that comes through the line 50 through the valve 51 and the line 52 into the steam distribution ring 53 that surrounds the vessel 22. '

If desired, further cooling by a water injection device (not shown) may be effected, which may preferably be directed to the zone of the connecting tubes between the cyclones.

Giant. 2 illustrates another embodiment of the invention that uses the side inlet of a separate used catalyst in physical association with a catalyst promoting oxidation from the reactor to the recovery vessels; The catalyst used enters the regeneration vessel 101 and flows downstream through the inlet duct 102 located on the side wall of the regeneration vessel 101 with a dense-phase orifice to the catalyst bed which is maintained at the lower section 106 at said orifice. at a small distance from the 10 * 7 phase boundary. Catalyst fluidization - is effected by the flue gas (air) entering via line 108 through outboard 109 and line 110 into air ring Ц1. Other air rings (not shown here) may be used to create an additional balance in the air flow through the regeneration zones. As described in Fig. 1, combustion of coke on catalyst particles within the dense phase zone begins where higher temperatures, if desired, can be achieved by transiently burning a heavy oil stream within the zone. Such oil can be fed by knowing -112 through the outlet 113 and through the conduit 114 terminating in the nozzle.

The air velocity of the fluidization can be controlled so as to continuously carry the catalyst particles upwards to achieve - heat absorption - in the dilute phase zone that forms the upper section 115 of the recovery vessel 11 > 1. that is, the section above the phase boundary - 10 * 7. Combustion of coke as well as carbon dioxide can continue in the dilute phase zone and to a large extent the utilized combustion products along with the entrained fraction of catalyst particles are discharged to the cyclone separators 120 and 121 of the first scab. Most of these catalyst particles are separated in first stage cyclones and discharged downwardly through tubes 122 and -123. submerged in the dense phase zone, into this phase. The gas and / or the remaining catalyst particles are then passed through the connecting lines 124 and 125 of the second stage cyclone separators 126 and 127, where in the undercoat the remaining amount remains.

The catalyst scavenging particles are separated and passed down through tubes 128 and 29. immersed in the bed of dense phase, into this phase. In fact. the spent combustion gas is then passed through line 130 and [31] to vessel [32] and finally removed from the recovery vessel 101 via line [33]. The regenerated dense bed catalyst is discharged via lines 134 and 135 equipped with collection heads 136 'and 137 for return to the catalytic cracker reactor.

As described in the embodiment of FIG. 1, carbon monoxide in the dilute-phase zone burns to form a high temperature zone in much of the dilute-phase zone, and in particular approximately in the area indicated by the letter X. by absorbing heat into the mass of the catalyst particles flowing upwardly by the exiting combustion gas stream. Temperatures in the region of the container 13'2. cyclones. and the connecting lines, may optionally be reduced by the supplied steam through the line 150 through the corrugated 151 and the lines 152 into the steam ring 153 which surrounds the container 132. Water spraying (not shown) may also be used - in a similar manner.

In another particularly preferred embodiment of the invention, the apparatus of FIG. 2 can be used with significant variations in operating parameters as compared to the embodiment described above. In this embodiment, the gas velocity and catalyst particle feed are adjusted such that the complete combustion of the coke and carbon monoxide within the dense phase is substantially completed and heat is dissipated in the catalyst bed. The stabilization of the combustion reaction is particularly enhanced by the use of an oxidation-promoting catalyst and the regeneration device can then be operated at lower temperatures, or - by burning a larger amount of carbon monoxide and thereby regenerating a larger amount of catalyst at the same temperature. .

Suitable hydrocarbon feedstocks for cracking processes include various mineral oil fractions with a higher boiling point than that of gasoline for use in engines, such as light oils, heavy oils, mixed gas oils, vacuum gas oils, gas oils - with a wide range fractions, kerosene, cracking residue, reduced petroleum fractions and oil fractions derived therefrom, as well as fractions obtained from bituminous shale oil, bituminous sands processing, synthetic oils, coal hydrogenation oils and the like. Such fractions may be used in one olive or in any desired combination.

The process according to the invention is characterized in that a substantial amount of coke and carbon dioxide is burnt in the dense phase at a substantially increased amount of catalyst particles compared to combustion in the dilute phase zone, which favorably dissipates the heat generated by the combustion. Since the portion of the combustion process occurring in the dense phase zone increases, the heat evolution in the dilute phase zone is substantially reduced and the need for rapid catalyst movement in the dilute phase zone to absorb the heat generated is reduced or eliminated at all.

_: Zvváště is desirable pouužtí methods of the invention identify an integral part of the device - for the cracking of a fluid cracking catalyst, as described above, the reactors are transmitted or vynášeným content while measures for separating zakoksovaných used catalyst, followed by regeneration of the catalysts used according to the invention.

Preferably, the cracking is carried out exclusively in reactors which utilize the catalyst suspension, without using a dense catalyst bed for cracking. Typically, where a fluidized catalyst cracking is used, there is a case of gas oil conversion, where the power ratio or volume ratio of total feed to fresh feed may be between 1: 2.

The conversion rate can vary from 40 to about 100 hmoo. % and preferably is maintained above about 60 wt.%, for example between about 60 and 90 wt.%. Co-conversion is herein understood to mean the percent reduction in the mass immo- bility of hydrocarbons boiling above approximately 220 ° C at atmospheric pressure to form light products or coke. - The catalyst to oil mass ratio of the fluidized bed reactor may be within the range of about 2 to 10 such that the fluidized dispersion will have a range of about 0/45 to 2.25 to 0.0283. the ratio between the catalyst and oil <m ee úrovni.ne maintained at greater than the above upper limit as výtotou in the range of ^{d 1,} 35 ^{d 2.2} 5 kg per ° ⁰²⁸ 3 m ^. The rate of fluorescence in ^a reactor of this kind may be between 6 and 18 m per second.

A reactor of this species to be 'preferably beneath Tata vertical, wherein the ratio of length to average diameter is at least about 25. Pro'výrobu - typical petroleum products ee keeps warm and ^t' and ^P 'at ODU reattoru ráctaání ^p řibH.žně 540 C ^to achieve C ^{u H} Inc evaporation of the raw material charge, so that in the upper section of the outlet temperature will be about 510 ° C.

Under these conditions, including measures to rapidly separate the catalyst used from the oil vapor outlet, a very short contact time between the catalyst and the oil is achieved. The contact time in the reactor of this kind is usually in the range between 3 to 10 seconds and preferably in the range between 3 to 7 seconds. Shorter contact times are - advantageous, because most. .

In the course of the cracking reaction, the initial contact time increases and the unwanted side reactions are avoided. This is particularly difficult when it comes to greater yield and selectivity, including less coke production.

Short contact between catalyst particles and oil vapors can be achieved by various means. For example, catalysts may be injected at one or more locations downstream of the reactor. Similarly, oil may be injected over the entire length of the lower section of the reactor and different injection sites may be used for the fresh stream - feedstock or feedstock recirculated. For example, the bottom section of such a reactor may comprise about 80% of the full length of the reactor in order to achieve extremely short effective contact times leading to optimal conversion of the oil charge.

Where another dense catalyst bed is used, provision should be made for catalyst particles and / or petroleum oviny feedstock to be injected directly into the dense bed zone. Although the conversion conditions reported herein are directed to the production of gasoline for internal combustion engines with an electric spark, the workflow may be varied appropriately to allow for the production of heavier hydrocarbon products such as fuel, jet engines, fuel for diesel engines and fuel heaters. oils.

Preferably, the spent catalyst from the crude oil conversion reactor is separated prior to entering the recovery vessel. The purging vessel used in the fluidized bed catalytic cracking apparatus may suitably be maintained at a temperature in the conversion reactor within the range of 457 to 570 ° C, and preferably about -510 ° C.

The preferred gas for the separation of the catalyst is steam, although nitrogen .You fed, other inert pl ^y n or ^p l ^y nn ^{e p} ro ^d uct ^s ^ s ^ Nominal For tlat typically ^R ° ^{from the} bounds on ^{d 0,} 7 kp cm @ ² to 2.45 kp / cm @ ² , which are suitable conditions for substantially removing all volatile components from the conversion catalyst used.

The dehumidified catalyst particles can be fed into the dense bed zone in the recovery vessel through the respective lines and through the valves of the separation vessel. The entry point may be at the bottom or side of the recovery vessel, preferably if the entry point is near the top of the dense bed of the fluidized bed. The feed can also be provided at the top of the recovery vessel, where the catalyst is contacted for the first time with the recovery gas used in a limited dilute zone.

The catalyst particles together with the oxidation promoting catalyst within the dilute phase may be partially entrained in the separation zone, typically containing cyclone separators in several stages, in which the catalyst may be returned directly by conduits immersed in the dense phase zone and the same. the regeneration and combustion gases are collected in a pressurized vessel and then vented to heat the energy contained therein in a suitable manner. The recovery of heat from flue gas includes steam generation, decontamination of spent catalysts, indirect heat exchange with various refining streams of liquids - and in particular with the feedstock to be converted, or use in various drying and evaporating arrangements.

When using a system arranged according to the first two embodiments described above, the recovery of the heat released essentially by the full combustion of coke and carbon monoxide is effected by absorption into the catalyser particles in both phases while returning the catalyst to the dense phase. dense phase band. The returned catalyst particles can bring additional heat to help raise the temperature in the dense phase to a temperature that favorably affects further removal of coke coatings on the catalyst particles so that substantially complete combustion of the coke deposit additive is achieved.

If the system is operated in such a way that substantially all of the burning is complete within the dense catalyst phase and the heat dissipates in that phase. when absorbed into the fluidized-bed particles, all the coke deposits are burnt. In all embodiments, the recovered catalyst returning from the recovery apparatus back to the cracking reactor suitably comprises from about 0.01 to 0.10 tip. %, particularly preferably from 0.01 to 0.05 wt. % and still about 0.01. to 0.03 hmoo. % carbon or coke, and may be removed from the recovery plant at a preferred temperature for use in a cracked reactor.

The regenerated catalyst particles having an unusually low residual coke content are recovered. recovered from the dense phase and defective at VP ^ <^ e ^ hundred ^<thee-by ^ C ^ o ^ ic ^ e ^ ající dense phase line to the cracking reactor in order to bring into contact with fresh hydrocarbon feedstock or mixtures thereof with hydrocarbon fractions re-circulated. Since the catalytic oxidation of carbon monoxide produced by the combustion of coke coatings on the catalyst can occur to a greater extent in the dense phase, and in preferred embodiments necessarily this dense phase oxidation is achieved, the recovered catalyst can return to the cracking reactor at much higher temperature. higher activity than conventional processes.

A number of fluid catalyst cracking devices are operated on the principle of thermal equilibrium, dependent on the combustion of the scythe to generate the heat required in the process. However, such devices have not been able to fully exploit the advantages of a cracking catalyst, especially zeolium catalysts, which can be particularly achieved in high-pressure catalysts where the contact time between the catalyst and the oil vapor can be extremely short.

The type of operation which achieves a high degree of conversion together with a high selectivity favorably affects the low ratio of catalyst potential to the amount of oil in a reactor of this type, resulting in less coke available to generate heat by combustion in the recovery plant. . Therefore, an additional external heat source, such as a feedstock preheater, must be provided to raise the temperature of the catalyst or to allow the device to operate at a lower temperature of the fresh feed.

These undesirable phenomena are possible or to a minimum extent by the process according to the invention which allows efficient recovery of additional heat from the regenerated catalyst particles for transfer to the reactor. The calorific value of coke in conventional operations is about 3,000 koalas per 0.45 kg.

The process of the present invention can still heat the heat available to coke at about 4.250 kcal per 0.45 kg. This higher combustion heat tends to increase the temperature of the regenerator, reduce the scrub rate on the regenerative catalyst, and reduce the catalyst circulation rate while maintaining improved yields at a given conversion rate.

A further advantage of the regeneration process according to the invention is that the recovered exhaust gas from the regenerator may contain an unusually low carbon monoxide content. While the flue gas during conventional cracking catalyst regeneration contains about 6 to 10 vol% carbon monoxide and similar propagation! carbon dioxide and very small quantities! oxygen,

The flue gas during regeneration according to the invention usually contains less than about 0.2% and often not more than 5 to 500 parts per million of carbon monoxide. The donkey content of the flue gas is not essential for environmental protection and may range from about 0.1% to about 10%, preferably from about 1% to about 3%, and more preferably;

in an amount not exceeding 2% to reduce the flue gas volume and maintain heat within the recovery reactor system.

If desired, any amount of carbon monoxide remaining may be appropriately burnt in the exhaust from the chimney exhausting the exhaust gas from the regenerator. From the point of view of operation, the method according to the invention can eliminate the method of recovering heat by burning carbon monoxide in plants or other processing arrangements, resulting in substantial savings on the plant and operating costs, while respecting the standards currently in force. ambient air quality in terms of carbon monoxide release.

the examples illustrate the process of the invention. Further, parts and percentages are by weight unless otherwise indicated.

He did

200 g of calcined conventional cracking catalyst, containing 5.3% hydrogen and rare earths with ion exchange, crystalline Y-type aluminosilicate, and silica and alumina blends with a total of 30% alumina are not lined with 3.90 g of 50% oanoan nitrate solution; 210 ml of water. About 80% of the catalyst is in dimensions. 20-75 µm · Impregnated catalyst particles are recovered and dried at 121 ° C, followed by calcination at 670 ° C for 3 hours. The resulting catalyst has a 0.3% manganese coating.

Pi a d 2

Example 1 was repeated by impregnating 1.265 g uranyl nitrate in 210 ml water. The impregnated catalyst is dried at 121 ° C and then calcined at 650 ° C for 3 hours. The catalyst has a content of 0.3% uranium.

She said d 3

Example 1 was repeated except that 0.82 g of tungsten tartrate dissolved in 210 L of water was used as the opaque solution. The canister was dried at 121 ° C and then calcined at 650 ° C for 3 hours. The resulting catalyst contains 0.3 wt% tungsten.

Example 4

Example 1 was repeated except that 2.39 g of ceroammonium nitrate dissolved in 210 ml of water was used as the impregnation solution. The catalyst was dried and calcined as in Example 1 and was found to contain 0.3% cerium.

Example5

Example 1 was repeated except that 2.73 g of zinc nitrate hydrate in a solution of 200 ml of water was used as the drinkable solution. The catalyst was dried and cured as in Example 1 and found to contain ~ 0.3 hmoo. % zinc.

Example 6

Example 1 - se. repeats, however, using 4.39 g of ferric nitrate in solution

200 ml - water as a solution. The impregnated catalyst was dried and calcined as in Example 1 and found to contain 0.3% iron.

Example 7

Example 1 is repeated and armored to be used. 1.1 g of ammonium molybdate in aqueous solution in 210 ml of water. The catalyst ae was dried at 121 ° C for 3 hours, then calcined at 650 ° C. The resulting catalyst contains 0.3 wt. % molybdenum.

P Hk lad!

Example 1 was repeated using 1.4 g of bismuth nitrate dissolved in 200 ml of dilute duuitic acid as the swelling solution. A dilute duuitic acid solution is prepared by adding 5 m concentrated. nitric acid to 1.4 g of bismuth nitrate and then dilute the solution with water to adjust the volume. The catalyst was dried at 121 ° C and then calcined at 650 ° C for 3 hours. The content catalyst is 0.3 wt. % bismuth.

Example. 9 ‘

Example 1 was repeated with the use of 5.0 g of titanium sulphate as an appealing solution. dissolved in 25 ml of aqueous 30% hydrogen peroxide, which is diluted to a volume of 200 µl with water. The solution is heated until the titanium salt is completely dissolved. The catalyst is dried at 121 ° C and then calcined for 3 hours at 650 ° C. The resulting catalyst contains 0.3 wt. % titanium.

Example 10

Example 1 was repeated except that 1.2 g of chromium trioxide dissolved in 200 ml of water was used as a solution for impregnation. The impregnated catalyst was dried at 121 ° C for 3 hours and then calcined at 650 ° C for 3 hours. The resulting catalyst has a coating of 0.6 wt. % chromium.

Prikladli ♦ *

Example 1 was repeated with 2.12 g of zigkunyl chloride dissolved in 200 ml of water as the impregnating solution. The impregnated catalyst was dried at 121 ° C for 3 hours and then calcined at 650 ° C for 3 hours. The resulting catalyst was 0.3 hmoo. % zirconium coating.

Example 2

Example 1.se. It was repeated with the use of 0.2596 g of a 50% manganese nitrate solution and 200 m of water as the drinkable solution. The impregnated catalyst was dried at 121 ° C for 3 hours and then calcined at 650 ° C for 3 hours. The resulting catalyst had a coating of 0.02 wt. % maanjanu.

Example 13

Example 12 was repeated except that 1.253 g of a 50% manganese nitrate solution in 210 ml of water was used as the latter solution. The resulting catalyst has a coating of 0.1 wt. % manganese.

Example 14

According to this example, a weight-type regeneration apparatus comprising a 3.49 cm inner glass unit in the fluidization section is used. 15.24 cm diameter as well as an outlet unit having an inner diameter of 4.3 cm by 12.7 cm. Synthetic mixture of gaseous flue gases contains! 4% by volume of carbon monoxide, 4% oxygen, 4% water vapor and 8Θ% nitrogen are passed through the fluidized bed catalyst with surface acceleration! gas 0.06 m per second.

The regeneration unit is surrounded by an oven to maintain the temperature at the desired level. The temperature of the flute bed is measured by thermoelectric cells. A first stage cyclone is arranged to separate entrained catalyst particles _mohly that could exit the recovery unit. The second stage cyclone is arranged downstream of the first stage cyclone to separate further catalyst particles that could be entrained by the exiting gas.

The catalyst separated in the first stage cyclone is fed back into the catalyst bed using a pipe leading into the bed. The off-gas from the second scab is further filtered using glass wool and analyzed by gas chromatography for oxygen, nitrogen, carbon monoxide and carbon dioxide.

The working time of the regeneration unit under the given conditions ranges from at least 40 minutes to about 90 minutes. This residence time is sufficient to give an oxidation state of critical importance to a particular oxidation catalyst at a given stage of the fluidized bed cracker.

In this example, the catalysts of Examples 1, 4, 5, 6 and 11 were used individually in this regeneration unit at an operating temperature of 650 ° C to determine the volume percent conversion of the carbon monoxide induced by the catalyst. The results are shown in Table I below.

Tabulkal

Oxidizing agent Weight % of total catalyst Order no. % conversion of carbon monoxide manganese 0.3 65 cer 0.3 72 zinc 0.3 55 iron 0.3 75 zircon 0.3 65 in the absence of this substances —- 31

Some of the catalysts mentioned above were tested using industry standard test methods for microfluidized catalyst units to determine the desired selectivity in catalytic cracking.

As the base, the catalyst without the oxidation promoting metal component has a relative microactivity of 154, a coke factor of 1.0, and a molar percentage of hydrogen to methane of 0.64. The catalyst of Example 1, when compared, shows a microactivity of 147, a coke factor of 1.1 and a molar ratio of hydrogen to methane of about 1.1 to 1.2. The catalyst of Example 4 shows a relative microactivity of 150, a coke factor of 1.1 and a molar ppocent ratio of hydrogen to methane of 0.9 to 1.1. The catalyst of Example 6 was found to have a relative activity of 134, a coke factor of 2.0, and a molar ratio of hydrogen to methane of 6.5.

Example 15

Example 14 was repeated using several catalysts containing 0.3 wt. % of the metal at different temperatures to determine the carbon monoxide convector s at these temperatures. The catalysts used are those of Example 1 and of manganese, of Example 4 and of Tsar, of Example 6e with iron and of Example 11 with zirconium. The results are set out in Table II.

225651

Table II

Oxidizing substance

The amount of impregnated metal in wt. %

Conversion of carbon monoxide by volume

595 ° C. 650 ° C. 685 ° C

manganese 0.3 32 - 50 cer 0.3 31 45 - iron 0.3 38 45 ' 55 zircon 0.3 - - 50 without the presence of this substance - 15 Dec 25 40 no catalyst in the bed - - 20 May 50-60

Examples 16 and 16

Example 14 was repeated using the same regeneration temperatures of 650 ° C, except that powdered metal oxide was used as the oxidation promoting component in the mixture with the cracking catalyst of Example 1. The results are shown in Table III.

Table III

Metallic oxide as oxidant Hmoo. % of total catalyst Order no. % conversion of carbon monoxide msaingančitý oxide 1.0 46 nitrous oxide 2,0 51 iron oxide 0.3 about 34 iron oxide 1.0 about 35 without the addition of this substance - 28

The powdered oxides added in this example have a diameter of approximately 5 5 or less, for example iron oxides have an average particle size of less than 1 µm.

Example 17

Example 16 was repeated using new oxidation-resistant oxides at a bed temperature of 650 ° C in the following tones. The results are shown in Table IV.

Table IV. · __ ‘____ - -.

Added Powder Operation Oxygen, wt. %

Order no. % conversion of carbon monoxide

0.1 ⁰ 1

Type of oxide; ferric oxide

as an agent 1 33 91 iron lanterns (Inlanid) 1 33 about 33 (Wisconsin) 1 30 63 90 2 33 46 rare earth oxides (Davison) 1 33 50

Table IV - continued

Powdered oxygen added, well. %

Operation

Order no. % conversion of carbon monoxide

0.1

0.4 1

Kerr-Mc-Gee 1 33 65 Oat Lattice (reagent) 1 27 Mar: 41 51

Example 18

Example 17 ee was repeated, except that the reaction was carried out at a temperature of 625 ° C from the addition of the powdered metal oxides listed in Table V.

Table V

The oxygen conversion to the carbon monoxide powder in% vol. % - metal with and without additives

0.3

Ingredient: iron oxide

(reagent) 49 95 to 100 ferric oxide (reagent) 52 90 to 100 oxide m8tLfat (technical) 51 75 to 80 meaigium oxide (reagent) 58 about 70

Example 19

Gas oil, characterized by a value of 23.4 ° API, boiling between 340-575 ° C. is cracked in a fluid-bed transfer type reactor at an average temperature. cracking at. 508 ° C. The reactor power ratio (total feed mass / fresh feed mass) was 1.34 and 36,000 bbl feedstock was fed per day. Catalyst particles correspond to Example 1 impregnated with an oxidation catalyst and circulate at a rate of 19.6 t / min. The weight ratio of catalyst in oil in the cracking zone is 3.7.

The reactor effluent is fed to the separation zone and to the cyclonic separator. The hydrogenated products are removed and the catalyst used is passed down through the cyclone and passed from it to a separation zone maintained at a temperature of 510 ° C. . The catalyst is pumped through the steam to remove residual volatiles before regeneration. Separate catalysts containing 0.9 wt. % of coke on the catalyst is fed to the bottom of the recovery vessel shown in Fig. 2, where the catalyst is airborne. in a catalytic bed of the heated phase α is maintained at a temperature between 685 to 690 ° C (average temperature is 670 ° C) with the combustion of coke and optionally with combustion of the fuel oil supplied, if necessary.

The air mass is stabilized at a rate of approximately 130 kg per hour to deliver approximately 6.3 kg of air per 0.45 kg of coke-coated catalyst. The catalyst is entrained by a rising air flow and transferred to the dilute phase zone of the catalyst at the top of the recovery vessel approximately at a position corresponding to the dense bed phase boundary. Combustion of carbon monoxide is completed in the dilute phase at about 760 ° C. The gases and the entrained catalyst are fed from the dilute phase zone to a series of cyclone 8, with the catalyst being directly returned to the dense phase zone. The gas stream leaving the cyclone system is first fed to a pressurized vessel located on top of the recovery vessel and then discharged at a temperature of 680 ° C to return to the transfer reactor.

Analysis of the regenerated catalyst shows that the residual coke content is only 0.03 wt. The analysis of the outgoing gases shows that the carbon monoxide content is zero and that the oxygen content is 1.9 wt. %. The cracking conversion was 67.7 vol% feedstock. According to thermal equilibrium calculations, coke is burned at a rate of 9 · 315 kg per hour, releasing 4.600 kcal from 0.45 kg of coke. The amount of heat generated above 80% is absorbed into the regenerated catalyst and thus maintained within the cyclic cracking fluid system.

Example 20

The same feed system and feed conditions were used as in Example 19 using the recovery vessel of Figure 2, but the regeneration gas rate and particle flow rate can be adjusted to absorb heat from the necessarily complete combustion of coke and oxide. by the action of the catalyst in the dense phase in the vessel. A level of carbon monoxide of 8 ppm can be achieved by the process of the invention to achieve coke removal and heat recovery in a manner similar to that described in Example 19 above.

Claims (10)

SUBJECT OF THE INVENTION Cyclic fluid catalytic cracking process, wherein the hydrocarbons are cracked in a fluidized state by contact with catalyst particles in a reaction zone, the coke-deactivated cracking catalyst is separated from the product taken from the hydrocarbon cracking reaction zone and is devoid of volatile matter. the volatiles are passed into the regeneration zone and fluidized therein with a regenerative gas containing molecular oxygen to form a lower dense phase of catalyst particles and an upper dilute phase of catalyst particles, the coke is burnt from the catalyst particles by contact with the regeneration gas, and as combustion products generates carbon monoxide and carbon dioxide, the gaseous product and the regenerated catalyst are withdrawn separately from the regeneration zone, and the regenerated catalyst is recirculated to the reaction zone for Other uses in hydrocarbon cracking, characterized in that the coke and carbon monoxide are burned in the regeneration zone up to a content of not more than 0.10% by weight of residual coke in the regenerated catalyst and not more than 1.0% by volume of carbon monoxide in the gaseous product. Maintaining a regeneration temperature of 537.77 to 815.55 ° C, the regeneration gas being fed to the regeneration zone in a stream supplying an amount of molecular oxygen in excess of 0.1 to 25% of the stoichiometric amount required for complete conversion of coke to carbon dioxide and vapor. A carbon monoxide-promoting agent comprising at least one metal with an atomic number of at least 20 and selected from groups IB, IIB and II to VIII of the periodic table is added to the catalyst particles, the greater part of the heat generated by the combustion of carbon monoxide being absorbed The catalyst particles are maintained in the catalytic cracking process by reintroducing the heated regenerated catalyst particles into the reaction zone. 2. The process of claim 1, wherein the coke and carbon monoxide are burnt in a dense phase. 3. The process of claim 1, wherein the coke is combusted from the catalyst in the dense phase of combustion of the carbon monoxide, completes in the diluted phase and the temperature of the diluted phase is at least 27.77 [deg.] C higher than the temperature of the dense phase. 4. A process according to claim 3, wherein the catalyst particles are circulated from the dense phase to the diluted phase of the regeneration zone, and are dispersed in the thin zone, and the greater part of the heat released by the dilution combustion is absorbed by these particles. particles introduced into circulation. 5. The process of claim 3, wherein the temperature in the dense phase is maintained at between 6413.68 ° C and 704.44 ° C and the temperature in the diluted phase is maintained at between 676.66 ° C and 787, 77 ° E 6. The method of claim 1, wherein the carbon monoxide oxidizing agent is supported by a solid ceramic carrier. 7. The method of claim 1, wherein the carbon monoxide promoting agent is supported by a solid support, selected from the group consisting of silica, alumina and mixtures thereof. 8. The method of claim 1, wherein the agent promoting oxidation of carbon monoxide is. supported by a cracked catalyst. 9. The process of claim 1, wherein the carbon monoxide oxidizing agent and the cracking catalyst are brought into powdered form in the regeneration zone. 10. The process of claim 1 wherein the carbon monoxide oxidizing agent comprises one or more metals selected from the group consisting of platinum, palladium and rhodium. '