Method of cyclical fluid catalytic cracking

(541 Cyclic Fluid Catalytic Cracking Method

BACKGROUND OF THE INVENTION The present invention relates to a process for catalytic cracking with reduced leakage of harmful gases, and more particularly to a process for cyclic cracking in a moving or an animal phase. The process is suitable for sulfur-containing hydrocarbon mixtures and leads to a significant reduction in the carbon monoxide and sulfur oxides content of the waste gases from the recovery plant.

The cracking catalyst, which is relatively deactivated during the cracking of the hydrocarbons in the reaction zone by carbonaceous deposits, commonly referred to as coke, is continuously discharged from the reaction zone. The catalyst so worn is passed from the reaction zone to a stripping zone where carbonaceous deposits are washed out and separated from the catalyst, and the catalyst is then led to a regeneration zone where catalyst activity is recovered by removing residual carbon deposits by burning coke in an oxygen-containing gas stream. to form carbon monoxide and carbon dioxide. The hot, regenerated catalyst is then continuously fed back into the reactor and the cycle is repeated.

In catalytic cracking, there is a problem with the incomplete combustion of carbon monoxide to carbon dioxide in the regeneration zone, and thus a certain amount of carbon monoxide is leaking from the regeneration zone. In addition to undesirable air pollution, carbon monoxide tends to react with residual oxygen in the waste gases, causing combustion in the ducts and smoke outlets and damaging these devices with excessive temperatures.

When sulfur-containing crude oil in the form of organic compounds is then treated in cracked catalyst systems, the coke which is deposited on the catalysts contains sulfur. In the recovery of coke-deactivated catalyst particles, coke is burnt on the catalyst surface while sulfur is converted to sulfur dioxide along with a small proportion of sulfur trioxide and this mixture enters the gas discharged from the recovery zone. When cracking a high sulfur feed, sulfur oxides are often leaked in the order of 1,200 parts per million (ppm).

Standardized conditions with regard to air purity are established for the discharge of carbon monoxide and fly ash, and similar regulations 1 for other waste gases such as sulfur oxides and in particular sulfur dioxide can be expected in the short term. For this reason, much attention is currently being paid to the problem of reducing leakage

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2! 4 S 7 5 7 fly ash and various products of combustion in waste gases from different combustion processes and from regeneration zones of cracking plants. To reduce the release of harmful products, it is therefore necessary to select processes which are effective but do not reduce the activity or selectivity of the cracking catalyst. In addition, unwanted leakage of harmful gases must not be replaced by another problem, such as an increase in fly ash leakage or an increase in operating costs. In view of these requirements, the desirable reduction in sulfur oxides leakage in the oil cracker is achieved by using a cracking catalyst, which is modified to minimize leakage of both carbon monoxide and sulfur oxides, while catalyst efficiency, stability and abrasion resistance are reduced. normal cracking conditions in existing or even new crackers do not change.

While we usually avoid the use of metals in cracking catalysts and is thought to be problematic to crack metal-containing paint in the presence of cracking catalysts, South African patent 7924/72 and its corresponding U.S. patent 3,909,392 disclose the possibility of using combustion catalysts or promoters together with cracking catalysts within the regeneration zone in the presence of a metal rod, grid, mesh or sieve in the combustion zone, as well as metal compounds which can be set in motion, in particular powdered oxides of transition groups, such as iron oxide, manganese dioxide and rare earth oxides, which are added to the catalyst to be added or are limited to regeneration plants. Belgian Patent No. 826,266 discloses a method similar to the aforementioned US patent using a cracking catalyst physically coupled to a carbon monoxide accelerating catalyst and is a metal with an atomic number of at least 20 and suitable oxidation promoters include metals from the group IB, IIB and III to VIII of the Periodic Table of the Elements, and in particular platinum, palladium, rhodium, molybdenum, tungsten, copper, chromium, nickel, manganese, cobalt, vanadium, iron, cerium, ytterbium and uranium. U.S. Pat. No. 3,808,121 discloses cracking recovery¹ a catalyst in the presence of a carbon monoxide oxidizing catalyst, the catalyst being used in the regeneration zone.

Belgian Patent 7,412,423 discloses a cracking catalyst comprising less than 100 parts per million, calculated per metal and recalculated further to the total catalyst, of at least one metal compound of Periods 5 and 6. This group is intended to lead to a particularly marked reduction in the carbon monoxide content of the flue gases by catalysis on the cracking catalyst. The patent also includes a molecular sieve catalyst prepared in the form of a sodium salt, exposed to ammonium ions during ion exchange, and then impregnated with rare earth metals.

As far as the leakage of sulfur oxides is concerned, various methods have been described for the treatment of waste and / or flue gases, such as scrubbing or scrubbing, chemical absorption, neutralization and chemical reactions or conversions, but all these processes for the removal of sulfur oxides , require bulky and expensive additional equipment, thus increasing both operating and investment costs. U.S. Pat. No. 3,699,037 discloses adding at least a stoichiometric amount of a calcium or magnesium compound, calculated on the amount of sulfur deposit on the catalyst, to the cracking cycle. This added material is to react with sulfur oxides to form substances which are then removed in a very fine state from the cracking cycle as a dusty material from the off-gas from the recovery plant. However, the continuous addition of the above substances clearly increases the operating costs. Similarly, U.S. Pat. Nos. 3,030,300 and 3,030,314 disclose a catalytic cracking process in which a single compound is continuously added to a moving-bed cracking cycle cycle, or more boron, alkali and alkaline earth metal compounds are added to prepare catalyst particles. with increased impact resistance as well as abrasion resistance of the surface, being silicon-containing catalyst particles, optionally silicon dioxide with a microporous catalytically activated core that is provided with an adhered protective glass coating consisting of silica and one or more boron, alkali or alkaline earth metal compounds.

U.S. Pat. No. 3,835,031 discloses a cyclic process of catalytic cracking in a liquid or animal phase, whereby a reduced leakage of sulfur oxides in the waste gases from the recovery plant is achieved. In the process, a catalyst is used which is a molecular sieve with a matrix, optionally with a support of silica and alumina, and the sieve is impregnated with one or more Group IIA metal oxides. U.S. Patents 3,388,077, 3,409,390 and 3,849,343 relate to a method for treating a stream of harmful off-gases containing carbon monoxide and sulfur oxides. The method comprises passing the stream through a catalyst system comprising a porous refractory support material, a catalytically active metal component, for example a platinum group metal, and an alkaline earth metal group from calcium, strontium and barium.

A new procedure has now been found.

The present invention relates to a cyclic method of ka. talytic cracking in the liquid or vapor phase, whereby a reduced release of sulfur oxides in the waste gases from the recovery zone is achieved. In various embodiments, the invention also results in a substantially complete combustion of the carbon monoxide in the recovery zone and in the absorption of heat released during such combustion on solid particles that are fed to the reaction zone and the stripping zone before being returned to the recovery zone. The solid particles are a molecular sieve cracking catalyst and a metal reactant; it can also be an amorphous cracking catalyst and solids which are substantially inert to hydrocarbon cracking.

The metal reactant can be incorporated into the cracking, molecular sieve catalyst, the amorphous cracking catalyst, and the substantially inert solid. And such incorporation can take place before or after introduction of the particulate substrate into the cracking process cycle. The cracking process cycle uses conditions such that a stable metal and sulfur compound in solid particles is formed in the regeneration zone, and the sulfur-containing gas is discharged from the stripper column.

SUMMARY OF THE INVENTION The present invention provides a process for cyclic fluidized catalytic cracking or catalytic cracking in the animal phase of hydrocarbon materials containing 0.2-6% by weight of sulfur in the form of organic compounds, in which the feed is cracked in the reaction zone with sieves, catalyst particles, the deactivated carbon and sulfur-containing deposits are separated from the reaction zone product and introduced into the stripping zone, where the volatile deposits from the deactivated catalyst are separated by stripping gas, the catalyst particles are then separated from the stripping product and transferred to the recovery zone and deposits containing sulfur and carbon, the oxygen-containing gas is recovered and the regenerated catalyst particles are separated from the regeneration zone gas and returned to the reaction zone, characterized in that the solid particles containing the molecular sieve-type cracking catalyst are allowed to cycle along with the metal reactant, the component comprises at least one metal in free form or in the form of a compound selected from sodium, scandium, titanium, antimony, copper, cadmium, rare earth metal, iron, chromium, molybdenum, manganese, cobalt, nickel, zinc and lead, the amount of the metal reactant being sufficient to absorb at least 50% of the sulfur oxides produced by the combustion of the sulfur and carbon containing deposits in the regeneration zone, the cracking being cracked in the presence of circulating particles temperature 454 to 649 ° C, the volatile separated products from the circulating particles are expelled at a temperature of 454 to 649 ° C with a stripping gas containing steam at a weight ratio of steam to a catalyst of catalyst 0.0005 to 0.025 per unit of time; 788 ° C, circulating! the particles absorb at least 50% of the sulfur oxides resulting from the combustion of deposits containing sulfur and carbon in the regeneration zone, Molecular oxygen-containing gas and traces of sulfur oxides and absorbed sulfur oxides as sulfur-containing material are removed from the recovery zone in volatile fractions from the reaction and / or ejection zone.

SUMMARY OF THE INVENTION The present invention provides an improved cyclic process for catalytic cracking in a liquid or animal phase by cracking a hydrocarbon feed containing organic sulfur compounds in a reaction zone using a liquid or animal layer using homogeneous or non-homogeneous recoverable moving solid particles. and the cracking catalyst is simultaneously deactivated by the action of sulfur-containing deposits. The entrained solid particles are separated from the cracked hydrocarbon stream from the reaction zone and fed to a stripping zone. wherein the deactivated catalyst is freed of separable hydrocarbon residues by contact with a suitable scrubbing gas. The movable solid particles are then separated from the gaseous stream from the stripping zone and fed to a regeneration zone where the treated deactivated cracking catalyst thus recovered is burned by incinerating non-ejectable sulfur-containing deposits and deposits by contacting the deactivated catalyst from the stripping column with an oxygen-containing stream to regenerate high activity, but to produce carbon monoxide, carbon dioxide, and sulfur oxides. The particulate solids containing the regenerated cracking catalyst are separated from the waste and / or flue gases in the regeneration zone and returned to the reaction zone.

The improvement of the process results in a reduced leakage of sulfur oxides in the waste gases from the regeneration zone, using solid particles comprising mainly a cracking catalyst of the molecular sieve type, which is a matrix or cracking catalyst substrate, crystalline aluminum silicates distributed over the matrix and metal a reactant with sulfur oxides to form a metal and sulfur containing compound, wherein the reaction takes place in solid particles; a steam-containing stripping gas is used, the deactivated cracking catalyst is recovered after the stripping operation at regeneration temperatures within the range of the metal-sulfur compound being stable, and a regeneration gas containing sufficient oxygen is introduced into the regeneration zone so that the gases discharged from the regeneration zone, contain molecular oxygen.

A suitable hydrocarbon feed in this process may contain 0.2 to 6% by weight of sulfur as organic compounds. Preferably, the feed comprises from 0.5 to 5% by weight of sulfur and in particular from 1 to 4% by weight of sulfur each in the form of organic compounds.

The molecular sieve-type cracking catalyst matrix is ​​preferably a combination of at least two materials selected from the group consisting of silica, alumina, zirconia, titanium, magnesium, thorium and boric oxide, and in particular the combination of silica with alumina. The cracking catalyst matrix preferably contains from about 10 to 65% by weight, in particular 25 to 60% by weight of alumina, preferably 35 to 90 and especially 35 to 70% by weight silica and 0.5 to 50 and especially 5 to 50% by weight crystalline aluminum silicate. The solid particles comprise a molecular sieve-type cracking catalyst generally 10 to 99.999% by weight, preferably 30 to 99.999%, and in particular 90 to 99.995%.

The metal reagent consists of at least one free or bonded metal element selected from the group consisting of sodium, scandium, titanium, iron, chromium, molybdenum, manganese, cobalt, nickel, antimony, copper, zinc, cadmium, rare earth metals and lead¹. Thus, a metal reagent may be used¹ select from sodium, scandium, titanium, iron; chromium, molybdenum, manganese, cobalt, nickel, antimony, copper, zinc, cadmium, rare earth metals, lead, their compounds and mixtures thereof. It is particularly preferred that the metal reagent consists¹ at least one free or bonded metal element from the group consisting of sodium, chromium, manganese, copper, zinc and cadmium. Ideally, the metal reagent is selected from at least one free or bound metal from the group of sodium, manganese, and copper.

It appears that the oxide or oxides of the metal elements or metal elements of the metal reagent are essentially responsible for the absorption of the sulfur oxides in the regeneration zone. For this reason, it is preferable to introduce the metal element (s) of the metal reagent into the catalytic cracking cycle in the form of the corresponding oxide (s). However, it is sufficient in the process according to the invention to add to the cycle at least one metal element suitable for use as a metal reagent. The metal element or metal elements of the metal reagent are activated to absorb sulfur oxides in the regeneration zone. This activation is likely to involve the partial or substantially complete conversion of the metals or metal of the metal reagent to the corresponding oxide or oxides. The activation is substantially independent of the manner in which the metal element is or is bound by the metal element at the time of introduction into the cycle.

The metal reagent is in the recovery zone in a moderate amount sufficient to absorb a major proportion of the sulfur oxides produced by the burning of deposits and deposits containing carbon and sulfur. Of the resulting sulfur oxides, at least 50% and preferably more than 80% is absorbed by the metal reagent in the recovery zone. Therefore, the concentration of sulfur oxides in the waste gases from the regeneration zone in the process of the invention is less than about 600 to 1000 ppmv, preferably less than 600 ppmv, and especially less than 400 ppmv.

The amount of metal reagent used, calculated on the metal or metals, is about 50 ppm to about 10% by weight, based on the total weight of the solid particles. With respect to copper, the mean amount, calculated as metal, is about 10 ppm to about 10%, in particular 50 ppm to 0.1%, and in particular 50 ppm to 250 ppm, based on solid particles. In the case of iron, this mean amount, calculated as iron, is about 50 ppm to about 5%, preferably 0.1 to 1.0%, and in particular 0.3 to 0.8%, based on the weight of the solid particles. In the case of the metal of the zinc, cadmium, manganese, scandium and cobalt groups, the median slackness, calculated as metal, is about 25 ppm to about 7%, in particular 0.01 to 5% and in particular 0.01 to 0.5%, based on mass of solid particles. When chromium, lead or antimony is used, their mean metal-based amount is about 10 ppm to about 1%, in particular 0.01 to 0.1%, and especially 0.01 why. up to 250 ppm, based on the weight of the solid particles. In the case of sodium, calculated on the metal, the mean amount is about 0.6 to 3%, in particular 0.8 to 2%, and in particular 0.85 to 1.5%, based on the weight of the solid particles. For titanium, the mean amount after conversion to metal is about 10 ppm to about 10%, in particular 0.5 to 1%, and in particular 0.5 to 0.8%, based on the weight of the solid particles. In the case of a rare earth metal, the mean amount, calculated on the basis of the rare earth, is about 0.2 to 10%, preferably 2 to 6%, and in particular 2 to 4%, based on¹ to the mass of solid particles. For nickel, the mean amount after conversion to metal is about 10 ppm to about 10%, in particular 50 ppm to 0.5%, and in particular 50 ppm to 0.1%, based on the weight of the solid particles. The above percentages are by weight.

The metal reagent is not separated on the solid particles¹ always evenly. Therefore, some particles may contain less or more than the abovementioned amount of metal reagent, but on average the solid particles contain the abovementioned amounts of metal reagent,

The deactivated catalyst, depleted in the stripping zone, is recovered in the stripping zone at temperatures to form stable metal and sulfur compounds from the metal in the solid particles and from the sulfur oxides. The regeneration temperatures are approximately 566-788 ° C and in particular 638-732 degrees Celsius. The hydrocarbon feed is cracked at reaction temperatures where the metal and sulfur-containing compound in solid particles react to form a metal sulfide in the metal reagent. The cracking temperature is generally about 454 to 649 ° C, and in particular 466 to 649 ° C. The expulsive deposits from the deactivated cracking catalyst are removed by treatment with a vapor-containing gas at temperatures where the metal sulfide in the metal reagent reacts with water to form hydrogen sulfide gas. The ejection temperatures are generally 554 to 649 degrees Celsius, and in particular 466 to 538 ° C. The weight ratio of steam to molecular sieve type cracking catalyst to be added to the stripping zone is about 0.0005 to 0.025 and preferably 0.0015 to 0.0225. The waste gases from the reaction zone preferably contain at least 0.1% by volume and particularly at least by volume 0.5% oxygen to achieve the desired reduction of harmful gases in the discharged waste gases.

According to one embodiment of the process according to the invention, the metal reagent is introduced into a cracking catalyst of the type. adds its sieve molecules. In this case, the metal reagent is introduced either into crystalline aluminum silicate or into a molecular sieve-type cracking catalyst matrix. According to another embodiment of the process according to the invention, the solid particles additionally comprise at least one material from the group of solids which are substantially inert in the cracking of hydrocarbons and also inert to the amorphous cracking catalyst, and the metal reagent is introduced into the material. According to a further variant of the process according to the invention, the metal reagent as such is a solid in solid particles.

The metal reagent can be incorporated into the solid particles either during the catalytic cracking cycle, in the cracking zone, in the stripping zone, in the regeneration zone, or outside these processes or zones. When incorporated during the catalytic cracking cycle, the metal reagent is added to the fluidized catalytic cracking cycle as an oil or water-soluble or dispersible metal or metal compound in the form of a solid, liquid or gas and can be introduced into solid particles in situ. Preferably, such a compound is selected from the group of ketone and metal compounds, metal carbonyl, metallocene, metal complexes with olefins of 2 to 20 carbon atoms, metal complexes with acetylenes, metal complexes and alkylphosphines or arylphosphines and metal salts with carboxylic acids of 1 to 20 carbon atoms. 20 carbon atoms.

The invention therefore relates to an improved process for catalytic cracking in a moving or an agitated phase, which is at the same time an improved process for regenerating the catalyst used in the catalytic conversion in a moving phase, and an improved process for reducing sulfur oxides in waste gases from cracking catalyst recovery zones. at the same time it concerns the conversion of sulfur-containing hydrocarbon feeds, wherein the cracking catalyst is deactivated by sulfur coke deposits on the catalyst surface. The solid particles used in the process of the invention include cracking catalysts of the molecular sieve type, circulate in a well dispersed state through a cracking process cycle that includes a cracking zone, a stripping zone, and a recovery zone. The conditions used result in a reduction in the sulfur oxide content of the waste gases from the recovery zone.

The cracking catalyst and the metal reagent of the invention have different functions. The cracking catalyst serves to catalyze the cracking reactions while the metal reagent is substantially inert with respect to the cracking reaction or has very little effect on the catalytic conversion reaction under the conditions used. To reduce the content of sulfur oxides in the waste gases from the recovery cycle, they absorb solid sulfur oxide particles in the recovery zone. The molecular sieve-type cracking catalyst often serves as such in the function of a sulfur oxide absorber. The metal reagent reacts with absorbed sulfur oxides to form a metal and sulfur containing compound, in particular to form a metal sulfate. Assuming that such a metal and sulfur containing compound is stable in the regeneration zone under operating conditions, it is transferred to the regeneration zone and the stripping zone on the solid particle surface, where it is reduced and separated as a sulfur-containing gas, particularly hydrogen sulfide.

The activity with respect to reducing the release of sulfur oxides in the waste gases from the recovery zone depends on the nature of the usable metal. With the usable metals in the metal reagent, no equivalent results are obtained as with the special metals which can be used in the metal reagent under different reaction conditions.

The solid particles according to the invention are finely divided and have a mean size of not more than 20 micrometers to about 150 micrometers, so that they are in a suitable form for moving or whirling. Suitable matrices for the cracking catalyst include silica and / or alumina. Other refractory metal oxides may also be used, the choice of which is ultimately limited by their ability to regenerate effectively under selected reaction conditions. Clays may also be used as an additive to alumina. Suitable catalysts are mixtures of silica and alumina mixed with molecular sieves, known as¹ zeolites or crystalline aluminum silicates. Suitable cracking catalysts are characterized by a content of crystalline aluminum silicates which is sufficient to increase the cracking activity of the catalyst, said amount being ultimately limited by the possibility of effective regeneration under the reaction conditions. Crystalline aluminum silicates usually contain silica and aluminum oxide in a molar ratio of at least about 2: 1, for example 2 to 12: 1 and preferably 4 to 6: 1. Crack catalysts on a silica support comprising, for example, 35 to 90% by weight of silica and 10 to 65% by weight of alumina are suitable. Such catalysts may be prepared using any suitable method, such as milling, gel co-precipitation, and the like, provided that the finished catalyst is in a suitable physical form for vortexing.

Suitable molecular sieves include both natural and synthetic aluminum silicate materials such as faujasite, chabazite, & quot; X & quot; and & quot; Y & quot; aluminum silicates, and are ultra-stable, or highly porous, crystalline aluminum silicates. When mixed with, for example, silica and alumina to form a catalyst for cracking oil, the molecular sieve content of the freshly prepared catalyst particles is suitably about 0.5 to about 50%, preferably 5 to 50% by weight. The equilibrium mixture of molecular sieve and cracking catalyst may contain only 1% of crystalline material. Crystalline aluminum silicates are usually swollen in sodium form or are prepared in this form. The sodium component is then reduced as much as possible, usually below about 0.3%, by ion exchange for hydrogen ions or precursors thereof, such as ammonium or polyvalent metal ions such as calcium, strontium, barium and rare earth metals such as they are cerium, lanthanum, neodymium and rare earth metals as found in nature, and such as mixtures thereof. The pore structure of the usable crystalline materials should be maintained under high temperature conditions during catalyst preparation, hydrocarbon processing, and catalyst regeneration. Crystalline aluminum silicates are often characterized by a uniform pore structure with a mean diameter of 0.6 to 2.0 nm and preferably of a mean diameter of 1.0 to 1.5 nm.

Catalytic cracking of heavy mineral oil fractions is one of the main refining processes for converting crude oil fractions into desirable flammable substances, such as high octane gasoline fuels and kerosene, used for internal combustion and spark ignition machines. One example of catalytic conversion is a method of catalytic cracking in the mobile or in the vapor phase whereby high molecular weight liquid hydrocarbons or their vapors are contacted with hot, solid and fine catalyst particles in either a moving reactor.¹ and in the ascending reactor, or in the extended ascending reactor, and the catalyst particle / hydrocarbon mixture is maintained at elevated temperature in a moving state or in an animal state for sufficient time to achieve the desired degree of cracking of the higher molecular weight hydrocarbon into the hydrocarbon fractions motor spirit and common spirits.

Suitable hydrocarbon feeds typically have a boiling point of 204 to 649 ° C when cracking, a higher range than the gasoline or kerosene fractions, and the cracking is generally carried out at 454 to 649 ° C. Suitable feedstocks include petroleum fractions above the boiling point of kerosene and hence gasoline, such as light gas oils, heavy gas oils, vacuum oils, kerosene, decanted oils, residual fractions, reduced crude oils and circulation oils derived from any from previous oils, as well as suitable fractions obtained from shale oils, from the treatment of tar sands, synthetic oils, coal liquefaction products and similar materials. These fractions may be used as such or in any suitable mixtures.

The process according to the invention can be carried out in any cracking apparatus and according to any scheme, but is preferably carried out in a catalytic moving or fluidized bed cracking system wherein the bulk of the hydrocarbon conversion takes place in the diluted phase in the transfer line or in the ascending reactor system. high active catalysts and relatively high flow rates are used. Preferably, the cracking takes place essentially exclusively in the upstream reactor and the subsequent dense cracking bed is not used for cracking. In a typical embodiment using cracking in an ascending reactor and gas oil conversion, the volumetric flow rate or volume ratio of the total feed to the fresh feed may be 1 to about 3. The degree of conversion may be about 40 to 100% by weight and preferably maintained above 60%, for example 60 to 90%. By conversion is meant the percentage reduction in the weight of hydrocarbons boiling above 221 ° C at atmospheric pressure to produce lower hydrocarbons or coke. The weight ratio of the total cracking catalyst to oil in the ascending reactor may be about 2 to 20 so that the dispersion that is moving or swirled has a density of about 16.1 to about 283 kg / m 2³. Preferably, the catalyst to oil ratio is maintained at about 3 to about 20, in particular 3 to 7. The velocity of movement or turbulence in the ascending reactor may be 3.088 to 30.48 m / s. The up-stream reactor is generally characterized by a ratio of length to mean diameter of about 25. When treating petroleum products, a mixed temperature in the bottom section of the up-stream reactor is preferably maintained at about 538-543 ° C to evaporate the oil feed; temperature 510 ° C. Essentially, higher temperatures are required for cracking residues and synthetic fuels. Under these conditions and assuming a rapid separation of the spent catalyst from the hydrocarbon vapor released, the contact time between the catalyst and the hydrocarbons is very short. The contact time in the ascending reactor is usually 1 to 15 seconds, preferably about 3 to 10 seconds. Preferably, short contact times are employed, since the highest proportion of hydrocarbon cracking occurs during the first contact times and thus avoids undesired secondary reactions. This is particularly important when higher product yields and higher selectivity are to be achieved as well as less coke formation.

Short contact times between catalyst particles and hydrocarbon vapors can be achieved in various ways. Thus, for example, the catalyst may be metered in at one or more locations along the length of the bottom or section at the bottom of the ascending reactor. In the same way, the hydrocarbon feed can be injected along the entire length of the bottom part of the upstream reactor, and different injection points can naturally be selected for both fresh stream and recycle streams. For this purpose, the lower part of the upstream reactor can be up to 80% of the total length of the upstream reactor in order to achieve an extremely short contact time, thereby optimizing the hydrocarbon feed conversion. If a dense catalyst bed is used, it is also possible to work with a device for directly feeding catalyst particles and / or hydrocarbon feed directly into the dense bed zone.

The above conditions are adapted to prepare fuels for internal combustion and spark ignition engines, but the process can also be adapted to maximize the production of heavy hydrocarbon products such as jet fuel, diesel oil, fuel oil and various chemicals, especially olefins and aromatics.

In the catalytic process, some of the non-volatile hydrocarbon materials or coke deposit on the catalyst particles. By coke is meant highly condensed aromatic hydrocarbons, usually containing less hydrogen, for example about 4 to 10% by weight. If the hydrocarbon feed contains sulfur in the form of organic sulfur compounds, the sulfur is also present in the coke. As the amount of carbon deposits on the catalyst increases, the cracking activity of the catalyst decreases as well as the selectivity of the catalyst in the preparation of gasoline blends. The catalyst particles can be regenerated again by removing a major portion of the coke from them by a suitable regeneration process.

The spent catalyst from the oil conversion reactor is stripped of expelled matter before it is admitted to the recovery plant. The stripping device used in conjunction with a moving bed cracking device can most effectively operate at the same temperature as it is maintained in the conversion reactor, i.e. at a temperature of about 454 to 649 ° C, and preferably at a temperature above 466 ° C. Steam is preferably used as the stripping gas, although steam-containing nitrogen or some other inert gas containing steam or waste gas or flue gas may be used. The purging gas is preferably used at an overpressure generally of about 0.68 to 2.41 MPa, which is suitable for substantially complete removal of the expellable compounds from the worn catalyst.

The method of the invention may be used in any method or in any recovery system after catalytic cracking, but is preferably used in a recovery system comprising at least one dense bed zone and one with a dilute phase. The worn catalyst particles after ejection of the ejection portions can be passed through a suitable conduit from the ejection device to the dense bed section of the recovery device. The entry may be through the bottom or the side and must be near the upper end of the dense bed moving zone. The inlet may also be directed to the upper end of the regeneration device, where the catalyst is first contacted with substantially spent regeneration gas in a limited zone of the diluted phase.

Catalyst regeneration is achieved by burning coke from the catalyst surface with a gas containing molecular oxygen, such as air. A number of regeneration processes are performed in which a high regeneration of catalyst activity is achieved depending on the coke removal rate. As coke is removed from the catalyst, the removal of residual coke is becoming increasingly difficult, and in practice, an intermediate stage of acceptable catalyst activity that can be achieved is chosen as an economic compromise.

Large amounts of oxygen or air are required to burn coke deposits on the catalyst. In a simplified form, coke oxidation can be regarded as carbon oxidation, according to the following reactions, which, however, do not imply any limitation: aj C + O2 -> CO2 bj 2 C -f- O2 -> 2 CO

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(c) 2 CC-j-Oz * 2 CO2

Reactions a) and b) both occur under typical catalyst regeneration conditions, where the catalyst temperature can be about 566 to about 788 ° C, and examples are known for the chemical action of oxygen on solid carbon and vice versa that occur when the catalyst is regenerated at temperatures within this range. The increase in temperature results in increased carbon combustion rate and improved removal of carbon or coke from the catalyst particles. Since the elevated combustion temperature is accompanied by a greater amount of heat generation, the gas phase reaction c) can also take place in the presence of sufficient free or molecular oxygen. The latter reaction is initiated by free radicals and can then be further controlled and catalyzed.

The combustion of the sulfur contained in the catalyst deposits on the catalyst also leads to the formation of sulfur oxides, which can be expressed by the following chemical equations, which, however, are not intended to be limiting.

d) S (in coke) 4- O2 -► SO2

e) SO2 + O22 O2 SO3

Reactions dj and e) also take place under typical cracking catalyst regeneration conditions. While reaction d) proceeds rapidly, reaction e) proceeds relatively slowly. However, reaction (e) can be catalyzed by the same catalysts that catalyze the previous reaction (c). Molecular sieves absorb oxides, sulfur, and reaction e) can therefore take place in a cracking catalyst on solid particles in accordance with the process of the present invention. Other proportions of solid particles can also absorb sulfur oxides. The resulting sulfur trioxide can then be reacted with a suitable metal or, in particular, a tin oxide in the metal reagent to form a stable metal sulfate in the solid particles. After the solids are separated from the waste gases of the recovery zone, the metal sulfate in the solids is returned to the reaction zone. Thus, sulfur leakage in the form of gaseous sulfur oxides in the waste gases from the recovery zone cannot occur.

The sulphate remains in the solid particles, which are returned to the cracking zone, where it is reduced to a metal sulfide corresponding to the metal in the metal reagent in a reducing environment and eventually passes to the sinusoid thanks. During the ejection with the steam-containing gas, sulphide is converted into hydrogen sulphide in the stripping zone, which is discharged into the waste gases from the stripping zone. In this way, the metal reagent is regenerated and restored to the reaction with the sulfur oxides in a further operation in the regeneration zone. Hydrogen sulfide can be recovered together with the cracked fractions from the ejector k-olons, separated and converted to elemental sulfur in the usual manner.

These reactions can be expressed by the following equations, but they do not mean any limitation:

regenerator:

M_x0 + SO2 + V2O2-> M_x0 + S0₃ - M_xSOd reactor:

M_xSO4 4- 4H2-M_XS 4 - 4 H2O - M_x0 4+ H2S 4- 3 H20 stripping device:

M_XS 4 - H2O - * M_x0 Wherein x represents the oxidation degree of the metal in the metal reagent upon reaction with oxygen.

These reactions are made possible by the molecular sieving catalysts and the presence of the metal reagent of the present invention. The high cracking activity, typically characterized by molecular sieve catalysts, is substantially unaffected by the presence of a metal reagent, so that the expected conversion of the feed occurs, and cracking products are formed in corresponding yields with reduced amounts of carbon monoxide and sulfur oxides leaking.

The metal reagent may be used in a very fine form, such as a powdered form, separately from the molecular sieve type catalyst or any other carrier. The metal reagent may be pulverized, mixed with a hydrocarbon cracking catalyst, and cyclically passed through the catalytic cracking process of the present invention. Generally, it is believed to be a preferred metal reagent so powdered that the powder is easily displaced and handled in a movable bed catalytic cracking system. The particle size of the powder should be chosen such that the particles do not divide beyond the bed speed.

Desirably, the particles of such a pulverulent product are not too small, thereby causing problems of excessive leakage of these particles with gases escaping from the bed and entraining such structured material. However, filters, cyclones, trapping devices, or separators, and the like are used to recover as much of the entrained particles as possible, so as not to lose them in progression. The powder should be strong enough to prevent excessive abrasion and further breakdown of the powder. Often, the mean particle size of the powdered metal reagent is about 0.5 to 100 microns, preferably 50 microns. It has been found that particles with a micro size, i.e. a mean sweat size of 1 micrometer, for example 0.01 to 0.5 micrometer, tend to agglomerate to form larger aggregates which are preferred for the process of the invention. Examples of powdered metal reactants which may be used are iron (III) oxide and iron (III) oxide, a mixture of iron (III) and zinc oxide, manganese (IV) oxide. cerium oxide and other oxides, as well as trimex, a substance described in U.S. Pat. No. 3,630,696.

The metal reagent may also be applied to a suitable carrier. The carrier can be an amorphous cracking catalyst or a solid which is substantially inert under cracking conditions and can be, for example, a ceramic material. In such a case, the metal reagent on the support is mixed with a molecular sieve catalyst. Preferably, the support is porous with a surface including pores of at least 10 m²/ gas preferably about 50 m²/G. Examples of suitable carriers are silica and alumina and mixtures thereof.

It is also possible to deposit the metal reagent of the cracking catalyst of the molecular sieve type or a portion of this reagent on the solid particles according to the invention. In this case, the metal reagent may be added to the cracking catalyst during catalyst preparation, or it may be impregnated with the cracking catalyst structure.

In any such case, careful consideration should be given to the choice of the method of incorporation, in order not to adversely affect the cracking activity and the selectivity of the cracking catalyst. If the cracking catalyst belongs to a type characterized by the presence of interchangeable ion sites, then it is preferred to carry out and complete the ion exchange prior to the addition of the metal reagent.

In all of the foregoing, the manner in which the metal or metals of the metal reagent is introduced into the cracking catalyst of the molecular, sieve, amorphous, catalyst or substantially inert substrate type is not known precisely. The metals come into complex combination with the carrier material and other solid particle components of the present invention. Therefore, it is understood that the use of the terms that a "metallic agent" is "introduced or incorporated" into the substrate includes metals as components which are found on the carrier in combined or bonded form and / or in an elementary state.

The metal reagent may be incorporated into the substrate by ion exchange, impregnation or other means, such as contacting the substrate or one component thereof with a solution or solutions of a compound or metal or metal compounds contained in the metal reagent in an appropriate amount to achieve the desired metal concentration. agents of the invention.

The metal reagent can be mixed with the substrate either at any stage in the preparation of the substrate or after the preparation of the substrate. One possibility of incorporating the metal reagent is ion exchange with the substrate. For example, it is preferable to ion exchange the crystalline aluminum silicate with a solution or solutions of the metal compound (s) or metal reagent (s), after which the product is combined with the porous cracking catalyst matrix after the ion exchange. It is also suitable to ion exchange solids containing silicon, optionally silicon dioxide or clay-like materials with a solution or solutions of a compound or compounds of metal or metals in a metal reagent. Suitable agents for this purpose are metal halides, preferably chlorides, nitrates, aminhalides, oxides, sulphates, phosphates and other water-soluble inorganic salts as well as carboxylic metal salts. acids with 1 to 5 carbon atoms, as well as the corresponding alcoholates.

Another process for preparing a metal reagent for use in the present invention is to impregnate a suitable carrier with a metal or metal metal compound, a water-soluble or water-dispersible agent, or an organic solvent. The impregnation can be carried out in any suitable manner which does not disrupt the structure of the substrate. The metal reagent can be impregnated with a hydrocarbon cracking inert carrier, a molecular sieve cracking catalyst, or an amorphous cracking catalyst.

Impregnation results in greater settling and primary physical association on the surface of the substrate surface, while the ion exchange leads to primary chemical association and greater diffusion and thus to reduced surface deposits. During impregnation, the metal is deposited and there is no substantial ion exchange between the metal and the substrate. When impregnating the substrate, the metal or metals of the metal reagent may be used in the form of a salt or salts soluble in water or in an organic solvent, dissolved in amounts sufficient for the substrate to contain the desired amount of metal or tusks. in contact with said solutions. The treated material is then freed from the solvent by drying, leaving the metal reagent deposited on the substrate.

The water-soluble nitric acid salts are preferably used for impregnation, since the residues after thermal decomposition of the nitric acid salts are relatively harmless with respect to the cracking catalyst activity in hydrocarbon cracking. Likewise, it is possible to use salts of hydrohalic acids or sulfuric acid for impregnation purposes, but because the by-products resulting from the thermal decomposition of these¹ These salts may be detrimental to the efficiency of the cracking catalyst, and are most often used when the metal reagent is applied to substrates which are substantially inert in the cracking reaction and do not adversely affect the cracking of the hydrocarbons.

Another possibility of physically depositing the co245757 reagent on a substrate, particularly on porous substrates such as crystalline aluminum silicates, is to adsorb the decomposable compound (s) in liquid form to the substrate, followed by thermal or chemical decomposition of the compound (s). The substrate can be activated by heating to remove any water present, thereby contacting the substrate with the degradable metal or metal compound of the metal reagent and thereby adsorbing the metal or metal compound to the substrate. Typical examples of such compounds include metal carbonyls, metal alkyl derivatives, volatile metal halides and the like. The adsorbable compound or adsorbed compounds may then be thermally or chemically reduced to the activated form, leaving the metal reagent uniformly dispersed on the substrate. The thermal reduction can be carried out, for example, in a regeneration plant during the regeneration process itself.

Both impregnation of the substrate and adsorption to the substrate can be carried out before the substrate is cycled. However, it is preferred to introduce the metal compound (s) of the metal reagent metal into the cracking process cycle and apply it to the substrate in situ. The appropriate compound (s) can be either in the form of an oil-soluble or dispersible material or, in water, or in solid, liquid or gaseous form, at any stage of the cracking process so as to achieve considerable dispersion over the solids. particles. Thus, for example, the compound or compounds used may be mixed with either the feed or swirl gas in the reaction zone, the regeneration gas, the torbanite or water in the reaction zone, or the stripping gas in the stripping zone, or the compounds may be introduced as separate current. Suitable in situ compounds include metal salts, organometallic compounds, metal diketone compounds, metal carbonyls, metallocenes, C 2 -C 20 olefin complexes, acetylene complexes, alkyl- or aryl-phenoic acid complexes, and salts. carboxylic acids having 1 to 20 carbon atoms, and, as special examples, the cyclopentadienyl sodium dicarbonyl dimer, dimethylzinc and diethylzinc dimer.

The activity and stability characteristics can be particularly easily achieved when the metal reagent is introduced into the cracking process cycle and applied to the solid particles in situ, as opposed to mixing or incorporating the catalyst during catalyst preparation. The addition of the metal reagent to the cracking process and its incorporation in situ results in a greater reduction in the amount of oxide leakage when compared to mixing with the cracking catalyst during its preparation. sulfur in waste gases from the recovery zone. The addition of the metal reagent during the cracking cycle is advantageous, and also because it achieves a higher effect with respect to the potentially harmful effect of such a metal reagent on the cracking reaction, further because the rate and / or amount of the metal reagent added can be varied. to cracker; cycle. There may also be a loss of the metal reagent which has been mixed with the cracking catalyst or has been introduced into the cracking cycle in the form of fine particles by abrasion of the cracking catalyst. Adding the metal reagent to the cracking cycle and incorporating the solid particles in situ allows the desired amount of the metal reagent to be maintained on the exterior surface or at accessible solid particle locations.

Advantageously, the invention operates using the recovery system disclosed in U.S. Pat. No. 3,909,392. It is directed to improve catalytic cracking, including an improved catalyst regeneration process for catalytic conversion of hydrocarbon feeds in a moving bed, wherein the catalyst is inactivated by coke deposits on the catalyst. surface of the catalyst. With the process of the present invention, the coke content of the regenerated catalyst can be extremely reduced, while maintaining a favorable thermal equilibrium in the conversion apparatus and producing a waste gas with an extremely low carbon monoxide content. The heat released by the combustion of carbon monoxide is adsorbed by the regenerated catalyst and forms part of the necessary heat in the hydrocarbon conversion zone. According to one embodiment of the process of the present invention, the combustion of carbon monoxide to carbon dioxide occurs substantially perfectly within the recovery plant in a relatively diluted secondary catalyst regeneration zone, preferably at a temperature of 649-816 ° C, particularly preferably at a temperature of 562-788 ° C. . The temperature in the secondary zone may be about 28 to 56 ° C higher than in the first regeneration zone. The partially regenerated catalyst from the relatively dense primary catalyst regeneration zone can be passed more rapidly through the secondary zone at such an amount and at a rate that it is substantially sufficient to adsorb all the heat released by combustion in the second zone. Although the largest portion of coke is burnt from the catalyst in the primary zone, another portion of coke from the partially regenerated catalyst is burned as the catalyst passes through the secondary zone, and a catalyst substantially free of coke can be recovered to recycle to the hydrocarbon conversion stage.

According to a second embodiment of the method of U.S. Pat. No. 3,909,392, substantially all combustion, including the oxidation of coke or carbon on the catalyst, and the oxidation of carbon monoxide in a single, relatively dense phase of regeneration takes place. zone due to favorable control mainly regeneration temperature and gas velocity. In the same manner, when carrying out the present embodiment of the present invention using the recovery system of U.S. Pat. No. 3,909,392, most of the heat released by the combustion of carbon monoxide in the recovery zone is adsorbed onto solid particles of the invention that include a cracking catalyst. to ensure no part of the heat required for the cracking zone. Such an embodiment of the process of the present invention allows for a substantial coke and carbon monoxide firing in a potentially dense phase zone, where substantially more solids are present than the dilute phase zone used, thereby dissipating the heat generated. As the proportion of combustion in the dense phase zone increases, the generation of heat in the thin phase zone is substantially reduced, thereby reducing or eliminating the need for rapid movement of solid particles in the diluted phase zone, which would be required to adsorb the heat generated.

In such embodiments, the process comprises the use of solid particles of the invention comprising a molecular sieve cracking catalyst and a metal reagent of the invention in a system where substantially complete combustion of carbon monoxide occurs. The low coke content on the catalyst thus obtained is below 0.2% by weight, preferably below 0.05%. Using this process, a carbon monoxide content of less than about 0.2%, for example about 500 to 1000 ppmv, up to a maximum of 500 ppmv can be achieved in the waste gases. The process also includes measures or devices for recovering the released heat by direct transfer to solid particles within the recovery plant.

In such an embodiment, the vortexing gas in the dense zone of the regeneration device may have a velocity of, for example, about 0.06 to 122 m / s, in particular 0.15 to 0.91 m / s. The regeneration gas used to move the dense band particles and having the function of the regeneration gas comprises free or molecular oxygen, and the oxygen is fed to the regeneration apparatus preferably in an amount such that it is slightly greater than that required to completely burn the coke carbon and hydrogen to carbon dioxide and steam. The excess oxygen in comparison with the amount just needed to completely burn the coke is 0.1 to about 25% or more with respect to the theoretical stoichiometric amount of oxygen; preferably, oxygen is used in excess of 10%. For example, when air is used as the regeneration gas, about 10% of its excess is about 2% of the excess oxygen in the off-gas. Preferably, the concentration of molecular oxygen or both free oxygen and carbon monoxide at each point of the recovery plant is kept outside the explosion limit under given conditions, and preferably the concentration of carbon monoxide is below the exploding carbon monoxide range to eliminate the risk of explosion.

The regeneration gas may contain, in addition to free or molecular oxygen, an inert gas or a diluent gas, such as nitrogen, steam, recycle gas from the regenerator waste stream. Often, the oxygen concentration in the regeneration gas when injected into the regeneration device is about 2 to 30% by volume, preferably about 5 to 25 percent. Since air is most effectively used as the oxygen source, nitrogen may be the major part of the inert gas. The inert gas can be used to distribute the excess heat from the coke combustion on the catalyst. The waste gas stream from the recovery plant is a source of hot inert gas and part of the gas can be returned to the recovery plant, for example together with a gas containing sufficient air or oxygen or possibly pure oxygen to achieve the desired oxygen concentration. Thus, it is possible to use recycle gas in direct heat exchange to raise the temperature of the regenerator gases to achieve even greater utilization of heat in the system.

The particulate solids in the diluted phase may be partially entrained into a separation zone comprising generally cyclonic separators in several stages, from where the solid particles are routed directly through the bottom line to the dense phase, and spent regeneration and combustion gases are collected in a reservoir; thereafter they can be discharged with appropriate recovery of the thermal energy contained therein. Methods for recovering heat from waste gases include steam generation, stripping of volatile components from spent catalysts, indirect heat exchange with various refinery streams, such as supply in some special processes, as well as use in various drying and evaporation equipment .

In the accompanying Figures 1 to 2, partial cross-sectional views are various embodiments of apparatuses suitable for regenerating the catalyst of the invention, including the regeneration scheme of U.S. Pat. No. 3,909,392. Indeed, such embodiments can be successfully used in a number of existing oil and hydrocarbon cracking plants, particularly in moving bed catalytic cracking with multiple space-intensive cracking, ejection and recovery arrangements.

Figure 1 shows one embodiment of the process of the present invention, wherein spent and washed catalyst from a non-depicted cracking reactor is passed through a bottom to a recovery plant. The solid particles containing the catalyst impregnated with the metal reagent from the injection zone together with the catalyst are discharged from the reactor and are discharged through the bottom into the recovery plant. These solid particles flow up through the inlet ducts 2 and 3 through the outlets 4 and 5 into the dense bed. The dense bed zone is at the bottom 6 of the regeneration device and occupies the space up to the boundary surface of the phases 7. The solid particles in the dense bed are set in motion or swirled by a stream of air to be incinerated and introduced through line 8, further through the closure 9 and duct 10 into the air ring 11. A substantially uniform flow of air in the regeneration zone can be achieved by using other air rings, not shown, as desired. Air combustion of coke on spent catalyst particles is initiated in a dense phase bed; higher temperatures can be achieved by temporarily burning torbanite oil, which is a bituminous shale with 70 to 80% carbon. Torbanite oil can be fed via line 12 through cap 13 and through line 14 which results in a nozzle positioned above the air ring 11. The air velocities causing the movement of the particles are continuously fed upwardly into the dilute phase zone occupied by the upper part 15 of the regeneration plant, i.e. the area above the boundary surface of phase 7. Coke combustion continues in the dilute zone and combustion gas consumed for the most part, along with the entrained solid particles, it is fed to the first stage of the cyclone separation devices 20 and 21. The major part of the solid particles is separated in the first cyclone stage and passed through the return lines 22 and 23 to the dense phase zone. The gases and the remaining solid particles are passed through cycles 24 and 25 between cyclones to the second cyclone stages 2G and 27, whereby substantially all of the remaining solid particles are separated, which are then led down through the submerged lines 28 and 29 into a dense band bed. The substantially consumed combustion gas is further passed through ducts 20 and 31 to space 32 and discharged from the regeneration apparatus via line 33. This discharged gaseous stream can be appropriately passed to a non-drawn heat exchanger, for example with a refinery stream, or used to produce steam. The solid particles containing the regenerated dense bed catalyst are passed through conduits 34 and 35, which are provided with collecting heads 36 and 37, back to the cracking reactor.

Although the combustion air supply usually leads to an excess of oxygen compared to the amount required to completely burn the coke on the catalyst particles to form steam and carbon dioxide, the coke may be combusted in a dense bed according to one embodiment of the present invention using a regenerative U.S. Pat. No. 3,909,392 is incomplete. In such embodiments, the combustion gases as they rise upward from the dense bed zone contain substantial amounts of carbon monoxide, as well as carbon dioxide and oxygen. The coke remaining on the catalyst, as well as the carbon monoxide, is substantially completely burned in the dilute phase zone with great heat evolution. If the carbon monoxide is burned in the dilute phase, the high temperature zone generally extends over most of the dilute phase zone, especially at the location marked with an "X", which can be easily observed by an unhindered window in this horizontal plane. The increase in the regeneration temperature within the diluted phase zone is partly influenced by the adsorption of heat by the mass of solid particles which are either entrained upward by a flow of rising combustion gases or escape via line 40 to the solids separation head 41 where the particulate fountain is divided into phase. The solid particles can be taken away by a stream of air, steam or other inert gas passing through a conduit 42 with a closure 43 and an orifice 44 leading to a small distance from the lower end of the tube 40. The excessively high temperature at the top of the recovery device can be further controlled steam, for example by means of conduits 45 and 46, a closure 47 and conduits 48 into the steam inlet opening 49. The temperature in the vicinity of the gas collection space can also be controlled by steam guided through the conduit 50 through the shutter 51, the conduit 52 and the steam ring 53 which is around the space 32. Optionally, further undisturbed water spray cooling can be provided. a conduit region 24 and 25 between the two cyclones. Lower temperatures favorably form the formation of stable metal and sulfur containing compounds in the regeneration zone.

Figure 2 is an example of another embodiment of the process of the present invention using the recovery systems disclosed in U.S. Patent 3,909,392, wherein the spent, devolatilized catalyst as well as the metal reagent coming from the reactor are fed to the recovery side · Equipment. The solid particles containing the spent catalyst impregnated with the metal reagent reach the regeneration device 101 via a line 102 which is very close to the boundary surface of the phases 107 downwards to allow entry into the dense phase bed which is The movement of the solid particles is caused by the air to be combusted, which is passed through the conduit 108 through the closure 109 and the conduit 110 to U.S. Pat. No. 3,363,993 (preheating of the feed). These heat recovery processes are designed to eliminate the discharge of carbon monoxide as part of the waste gases into the atmosphere and to avoid significant environmental burdens. Procedures require custom equipment.

By contrast, catalysts based on silica and alumina, as commonly used for many years in various petroleum hydrocarbon cracking processes, are not very sensitive to the residual coke content on the catalyst, provided that the coke content does not increase above 0, 5%. However, silica-based and alumina-based catalysts have been largely displaced by other catalysts which additionally contain a crystalline component of aluminum silicate and are known as zeolites or molecular sieves. Catalysts containing molecular sieves are substantially more sensitive to residual coke and are greatly influenced by both catalytic activity and catalyst selectivity when converting the hydrocarbon feed to the desired product or products. Taking into account the difficulties associated with conventional catalyst regeneration techniques in removing the last traces of residual carbon deposits, the coke content practically corresponds to the rest of the coke on the regenerated catalyst in the range of 0.2 to 0.3% by weight.

Since the increased activity and selectivity of cracking catalysts of the molecular sieve type is achieved with low coke content, there is a strong incentive to seek means to further reduce the remaining coke content. A coke content below about 0.05% by weight is particularly desirable, but is usually not achieved by conventional process equipment. Considerations of larger regeneration plants, more catalyst, greater heat loss, and the like are far from suited to solving catalyst problems in terms of ideal balance and activity.

Many moving bed crackers operate on the principle of thermal equilibrium, which is dependent on the combustion of coke to produce as much heat as is necessary for the process. However, it is not possible to fully exploit all the advantages of cracking catalysts, especially of the zeolite type, which can be achieved in a single ascending reactor, where the contact time between the catalyst and the oil vapors is extremely low. The high conversion process together with the high selectivity utilizes a favorably low catalyst to crude oil ratio in the upstream reactor, resulting in lower coke levels due to heat generation in the reactor combustion. As a result, an external heat source is often required to preheat the feed to raise the catalyst temperature, or else operate at a lower fresh feed temperature. Such undesirable process characteristics can be eliminated by the process of the present invention, or reduced to a minimum, allowing efficient recovery of additional heat by means of solid particles to be transferred to the upstream reactor. In a conventional coke incineration process, the heat generated is approximately 35,520 kj / kg of coke. By the process of the invention, the heat obtainable by the combustion of coke can be increased up to at least 37 200 kJ per kilogram of coke. This higher combustion heat affects the temperature in the recovery plant, the coke content of the recovered catalyst decreases, and the circulation velocity of the solid particles also decreases while achieving increased yields at a given degree of conversion.

In the following examples, parts are by weight unless otherwise indicated.

Example 1

200 g of a calcined, commercially available equilibrium cracking catalyst of the type consisting of a cracking catalyst mass, a crystalline aluminosilicate and a matrix containing 5,3% hydrogen treated by the exchange of rare earth ions obtained from crystalline silicate of aluminum Y-type and a mixture of silica and alumina containing 30% alumina are impregnated with 3.90 g of a 50% manganese nitrate solution and 210 ml of water. About 80% of the catalyst is 20 to 75 microns. The impregnated catalyst particles were dried at 121 ° C, calcined at 677 ° C for 3 hours; the catalyst then contains 0.3% manganese.

Example 2

The procedure of Example 1 is repeated, but a solution of 1.265 g of uranyl nitrate in 210 ml of water is used for impregnation. The impregnated catalyst was dried at 121 ° C and calcined at 649 ° C for 3 hours. Contains 0,3% uranium by weight. Example 3

The procedure of Example 1 was repeated, but a solution of 2.35 g of cerium ammonium nitrate in 200 ml of water was used for impregnation. The catalyst was dried and calcined as in Example 1 and contained 0.3 wt% cerium.

Example 4

The procedure of Example 1 is repeated, but a solution of 2.73 g of zinc nitrate x272757 in 200 ml of water is used for impregnation. After drying and calcining as in Example 1, the catalyst contained 0.3 weight percent zinc.

Example 5

The procedure of Example 1 was repeated, but 4.35 g of ferric nitrate in 200 ml of water was used as the impregnation solution. After drying and calcining as in Example 1, the catalyst contains 0.3 wt% iron.

Example 6

The procedure of Example 1 is repeated, but a solution of 1.1 g of ammonium molybdate in 210 ml of water is used for impregnation. The impregnated catalyst was dried at 121 ° C and then calcined at 649 ° C for 3 hours. The finished catalyst contains 0.3 weight percent molybdenum.

Example 7

The procedure of Example 1 is repeated, but a solution of 5.0 g of titanium sulfate in 25 ml of an aqueous 30% hydrogen peroxide solution is used for impregnation, and the solution is diluted to 200 ml with water. The solution is heated until the titanium salt is completely dissolved. The catalyst was dried at 121 ° C and then calcined at 649 ° C. The final catalyst contains 0.3% by weight of titanium. Example 8

The procedure of Example 1 is repeated, but a solution of 1.2 g of chromium trioxide in 200 ml of water is used for impregnation. The impregnated catalyst was dried at 121 ° C for 3 hours, then calcined at 649 ° C for a further 3 hours. The catalyst thus prepared contains 0.6% by weight of chromium.

Example 9

The procedure of Example 1 was repeated using 0.2506 g of a 50% manganese nitrate solution in 200 ml of water for impregnation. The impregnated catalyst was dried at 121 ° C for 3 hours and then calcined at 649 ° C for 3 hours. It contains 0.02% by weight of manganese.

Example 10

The procedure of Example 9 was repeated using 1.253 g of a 50% manganese nitrate solution in 210 ml of water for impregnation. The catalyst obtained contains 0.1% by weight of manganese.

Example 11

Dissolve 100 mg of chloroplatinic acid in a liter of water, and dilute 18 ml of the solution obtained with enough water to soak 300 g of the usual equilibrium cracking catalyst. The catalyst was previously calcined for 5 hours at 538 ° C and contained 2.5% molecular sieve by weight and 0.6% sodium. The wetted catalyst was dried at 121 for 3 hours^CC, calcined at 538 ° C for 3 hours, and then contains 6 ppm platinum, calculated on the total weight of the catalyst.

Example 12 g of conventional alumina are moistened with a solution of 3.2 g of ammonium vanadate and 5 g of oxalic acid in 95 ml of water, and after drying at 121 ° C for 3 hours, the catalyst is calcined at 538 ° C for 3 hours. The vanadium-impregnated catalyst is then moistened with a solution of 9.3 g of copper nitrate in 95 ml of water. The moistened alumina was dried at 121 ° C for 3 hours and calcined at 538 ° C for 3 hours. Aluminum oxide contains 2.5% by weight of vanadium and

2.5% copper.

Example 13 g of a solution of 6.9 g of a lubricating oil additive containing 9.2% by weight of magnesium (such as magnesium hydroxide, magnesium carbonate and magnesium salt of polypropylbenzenesulfonic acid in 33.1 g of light oil from the catalytic cycle) is run in an experimental cracking apparatus using a bed of 200 g of a conventional cracking catalyst containing 2.5% by weight molecular sieves and 0.6% sodium, is cracked, after which the solid is taken from the animal⁷ bed and calcined. The cyclized oil, the intermediate distillate obtained in the catalytic cracking, boiling at 221-343 ° C, was cracked at 371 ° C for 4 minutes. After purging the catalyst bed with nitrogen (10 minutes) at 677 ° C, the catalyst bed was cooled to 371 ° C and the cracking, purging and regeneration cycle was repeated until the catalyst reached a magnesium content of 1100 ppm, zinc 703 ppm and phosphorus 59 ppm. Example 14

The procedure of Example 13 is repeated but the cracking, flushing and regeneration cycle is carried out with 10 g of a solution of 6.5 parts of oil and 3.5 parts of lubricant additive containing 1.6% zinc, 1.3% phosphorus and

4.6% magnesium to a catalyst content of 2,400 ppm, 1,200 ppm zinc and 1,097 ppm phosphorus.

Other non-drawn rings may optionally be used to achieve an additional balance of air distribution over the regeneration zone. As described in connection with Figure 1, coke firing is initiated from the spent catalyst particles in the dense phase zone, where a higher temperature can possibly be achieved by temporarily burning the torbamate oil stream in the zone. This inlet is admitted by means of a duct 112, a closure 113 and a further duct 114 that opens through a nozzle.

With the speed of the air to be moved or the vortex of the particles, it is possible to control the continuous entrainment of the solid particles for heat absorption into the diluted phase which is in the upper part 115 of the regeneration plant, i.e. The carbon monoxide may further continue in the dilute phase zone, and the largely worn combustion gas, along with entrained solids, is removed in the first stage of the cyclone separator 120 and 121. The bulk of these solid particles are separated in the first cyclone stage, and are then passed through the dip lines 122 and 123 as a downstream stream into the dense phase zone.

The gases and remaining solid particles are then passed through lines 124 and 125 between. cyclones to the second stage of the cyclamic separators 126 and 127, where substantially all the remaining solid particles are separated, which are discharged by the plunger lines 128 and 129 downwardly into the dense phase bed. The substantially consumed combustion gas is then passed through lines 139 and 131 to the air space 132 and then discharged from the regeneration device via line 133. The solid particles containing the regenerated dense bed catalyst are discharged through tubes 134 and 135 provided with heads 136 and 137 and recycled to the reactor for catalytic cracking.

as described in connection with the embodiment of Figure 1, the carbon monoxide is burned in the diluted phase to form a high temperature zone, in large part of the dilute phase zone, and is particularly manifested at the location indicated by X. The dilute phase is largely achieved by the absorption of heat by the mass of solid particles supplied by the rising combustion gas. The temperatures around the chamber 132, the cyclones and the connecting ducts can optionally be lowered by steam passing through the duct 150 through the shutter 151 and then via a duct 152 to the steam ring 153, which extends around the area 132. water.

In yet another particularly preferred embodiment of the method of the present invention, the apparatus of Figure 2 is used to substantially alter the operating parameters in alignment with the embodiment just described. In this embodiment, the velocity and the injection of solid particles are adjusted so that the coke and carbon monoxide are substantially burned in the dense phase and heat is distributed throughout the bed.

When operating in the system according to one of the two embodiments just described, heat is released by substantially complete combustion of coke and carbon monoxide by adsorption on the solid particles in both phases and returning the solid particles to the solid phase, thereby also ensuring maintaining a stable high temperature in the dense phase band. The returned solid particles can bring with them additional heat, which serves to raise the temperature of the dense phase zone to a temperature that favorably affects the combustion of the additional crust of coke deposits so that the combustion of the remaining coke deposits occurs substantially completely. When the system is operated in such a way that substantially all combustion takes place entirely in the dense catalyst phase. then the heat is transferred by the solid particles in motion and the remaining coke deposits are burned. As a result, in all embodiments of the process, the solid particles containing the recovered catalyst, which are recycled from the regeneration apparatus to the cracking reactor, consist in a total weight of about 0.01 to about 0.10%, preferably 0.01 to about 0.05. % and in particular from about 0.01 to about 0.03% by weight of carbon or coke on the catalyst, and may be removed from the recovery plant at a preferred temperature for use in a cracking reactor.

An essential advantage of the process of the invention is to provide a regenerated catalyst with generally improved activity and selectivity that is very close to similar properties of a freshly prepared conversion catalyst, especially when used in conversions that are achieved at very short contact times in ascending reactors. Catalytic cracking activity. containing molecular sieves, as well as their selectivity in converting hydrocarbon feeds to the desired products, is always strongly and in a favorable sense improved by removing the remaining carbon or coke from the catalyst during regeneration. The lower coke content of the regenerated catalyst is particularly advantageous in the case of movable bed catalysts containing catalytically active crystalline aluminum silicates. Thus, higher yields of the desired conversion products can also be achieved.

In the cracking process using a dense phase lower zone and a diluted phase upper zone in the regeneration zone, oxidation of carbon monoxide to carbon dioxide in the dense phase of the regeneration plant can often be at least about 60% and often 65 to 95% to complete. Oxidation of carbon monoxide to carbon dioxide in the dense phase generates heat, which promotes uniform combustion of coke deposits on the moving bed catalyst. In addition, due to the oxidation of a substantial fraction of carbon monoxide in the dense phase, there remains less carbon monoxide to burn in the upper phase of the movable bed catalyst in the recovery plant, thereby substantially reducing or completely preventing "post-combustion" and high temperatures due to uncontrolled and over-combustion of carbon monoxide at the top of the recovery plant, which adversely affects reactor design, gas evacuation, equipment, intended to collect particulate materials from the off-gas, such as cyclones, and the phenomenon may also affect the activity of the catalyst.

The solid particles containing the regenerated catalyst with an unusually low coke content are thus recovered from the dense phase and passed into a cracking reactor via a conduit at essentially a dense bed temperature with fresh hydrocarbon feed or with a mixture of fresh hydrocarbon and hydrocarbon fractions which are recycled. Since the oxidation of carbon monoxide resulting from the combustion of coke deposits on the catalyst can largely take place in the dense phase, and according to a preferred embodiment of the process according to the invention, proceeds substantially completely in the dense phase, the recovered catalyst can be fed into the cracking reactor at substantially higher temperatures and higher activity than previously possible under conventional procedures.

The main advantage of the process according to the invention is in the unusually low carbon monoxide content of the waste gases from the recovery plant. If the off-gas usually contains about 6 to 10% of carbon monoxide, a similar amount of carbon dioxide and very little oxygen after the usual recovery of the cracking catalyst, a very low amount of carbon monoxide in the off-gas can be achieved using the regeneration process according to the invention. 0.2%, for example 500 to 1000 ppmv, calculated on the gas volume. Preferably, the carbon monoxide content is even lower, for example 0-500 ppmv. Such a low concentration of carbon monoxide in the waste or flue gas permits the direct discharge of the waste gas into the atmosphere, as well as the requirements of air purity. If desired, the remaining carbon monoxide can be burned in a suitable manner in the exhaust from the recovery plant. This advantage according to the invention makes it possible, inter alia, to eliminate the investment costs of installing carbon monoxide combustion plants and associated turbine-type plants or other partial energy recovery devices resulting from the subsequent oxidation of carbon monoxide, while still maintaining stringent requirements in terms of air pollution by carbon monoxide leakage.

The process according to the invention has further advantages. These relate to the problem of post-combustion and temperature equilibria. The main problem that often arises in practice, especially when regenerating moving bed catalysts, and which should be avoided, is the phenomenon referred to as · post-combustion as discussed, for example, by Hengstebeck in "Petrolem Processing", McGraw-Hill Co., 1959, pp. 160 and 175, hereinafter "Oil and Gas Journal", 53 (No. 3), 93-94 (1955). The term refers to the further combustion of carbon monoxide to carbon dioxide according to equation c) hereinabove, wherein the reaction is extremely exothermic. Post-combustion is extremely undesirable in the catalyst regeneration process because it is associated with very high temperatures which damage the equipment and can permanently deactivate the cracking catalyst particles. Numerous precautions in connection with the regeneration of movable catalysts cause additional combustion, and for the technology this condition is very difficult by describing a number of measures to control the regeneration process so as to avoid additional combustion. Recently, attempts have been made to raise temperatures in a recovery plant for a variety of reasons, and an arrangement has been described in which the temperature in the recovery plant can be maintained in the region of incoming post-combustion by appropriate measures to increase oxygen supply to the reactor. U.S. Patent Nos. 3,161,583, 3,206,393 and 3,513,087. According to the practice, the off-gas used in processes that prevent the post-combustion usually contains very little oxygen and a significant amount of carbon monoxide and carbon dioxide in almost equimolar amounts.

However, further combustion of carbon monoxide to carbon dioxide is an interesting source of thermal energy, since reaction c) is strongly exothermic. The post-combustion may take place at temperatures above about 593 ° C and represents about 10,674.4 kJ / kg of oxidized carbon monoxide. This represents typically about a quarter of the total heat released, calculated on the heat that is released when the coke is completely burned. The combustion of carbon monoxide can be carried out in a controlled manner in a separate carbon monoxide reactor after separation of the waste gases from the catalyst, as described, for example, in U.S. Patent 2,753,925, utilizing the amount of heat released for numerous refining processes such as high pressure steam. Other uses of this thermal energy are described in U.S. Patent Nos. 3,012,962 and 3,137,133 (turbine propulsion), and

Example 15

The procedure of Example 14 is repeated, but using a conventional equilibrium catalyst containing 3.3% molecular weight. sieves in a silica-alumina matrix taken from a conventional moving-bed catalytic cracker, calcined and the cracking, purging and regeneration process repeated until the magnesium content of the catalyst dropped to 4,600 ppm, zinc to 340 ppm and phosphorus to 1136 ppm.

Example 16

To determine the effectiveness of a certain number of impregnated catalysts of Examples 1 to 15 with a view to reducing the leakage of carbon monoxide from the recovery zone in the waste gases, a laboratory scale recovery apparatus is used. Synthetic off-gas or flue gas, each containing 4% by volume of carbon monoxide, 4% oxygen and 4% water vapor, 88% nitrogen, is fed at a rate of 1,000 ml (measured at 15.6 ° C) through a fixed moving bed. or a fluid bed of a molecular sieve-type cracking catalyst, wherein the catalyst is impregnated with metal, the catalyst being placed in a glass regeneration device connected to a heating system by which the temperature can be raised to a regeneration temperature of 649 ° C. The catalyst temperature is measured with a thermocouple. A cyclone is used to separate entrained particles from the gas, whereby the catalyst thus trapped from the escaping gas is returned to the catalyst bed. The operating time of the regeneration device under the given conditions is 40 to about 90 minutes in order to allow sufficient time to reach the metal oxidation degree on the catalyst in the effective catalytic cracking unit.

The gas escaping from the recovery plant is analyzed on a gas chromatograph for oxygen, nitrogen, carbon monoxide and carbon dioxide content. The amount of oxidized carbon monoxide is determined from the difference between the content of carbon monoxide in the fresh synthetic gas mixture and the content of carbon monoxide in the gas from the recovery plant.

In Examples 16-19, the impregnated catalyst of Examples 1, 3, 4 and 5 was used, and the volume percentages of oxidized carbon monoxide were 65, 72, 55, 75 and 65 respectively. In Example 20, the impregnated catalyst of Examples 1, 3, 4 and 5, and the conversion of carbon monoxide was 31% by volume.

Examples 21 to 24

Some of the prior art catalysts have been tested by standard test procedures as used in industry in a micro-movable bed catalysing apparatus with a view to determining the desired selectivity of the catalytic cracking process. As a basis in Example 21, the non-impregnated cracking catalyst used initially in Example 1 had a relative microactivity of 154, a coke factor of 1.0, and a ratio (in mole percent) of hydrogen to methane of 0.64. In Example 22, the impregnated catalyst of Example 1 was used, characterized by a relative microactivity of 1.47, a coke factor of 1.1, and a ratio (in mole percent) of hydrogen to methane of about 1.1 to 1.2. used in Example 3 and was characterized by a relative micreactivity of 150 coke factor of 1.1 and a hydrogen to methane ratio (in mole percent) of 0.9 to 1.1. In Example 24, the catalyst prepared in Example 5 was used having a relative activity of 134, a coke factor of 2.0, and a ratio (in mole percent) of hydrogen to methane of 6.5.

Examples 25 to 31

The procedure of Examples 16 to 20 is repeated, but powdered metal oxides having a particle size of not more than 5 microns or less are used in admixture with the impregnated catalyst of Example 1 instead of impregnated catalysts. Powdered metal oxides, their amounts and percentages by volume of carbon monoxide converted to carbon dioxide in Examples 25 to 28 as well as Example 29 for comparison when no metal oxide was used but the other conditions were identical are shown in Table 1. The same data for Example 30 and Comparative Example 31 where no metal oxide was used under otherwise identical conditions are also shown in Table 1.

Taboo

Example Metal Oxygen

<td> 25</td><td>manganese dioxide</td>

<td> 26</td><td>manganese dioxide</td>

<td> 27</td><td>iron oxide</td>

<td> 28</td><td>iron oxide</td>

<td> 29</td><td> —</td>

<td> 30</td><td>rare earth oxide</td>

<td> 31</td><td> —</td>

<td>Examples</td><td>32 and 33</td>

The procedure of Examples 16-20 is repeated, but a synthetic waste gas containing 1500 ppm of sulfur dioxide in a 4% v / v / v / v mixture of nitrogen is used; this gas is passed at 677 ° C through a regenerator at a rate of 1500 ml per minute (measured at

15.6 ° C). The waste gas is analyzed by means of an ultraviolet analyzer and the sulfur dioxide content is monitored continuously. Comparative experiments were carried out using the impregnated catalyst of Example 11 and the impregnated catalyst also of Example 11. In Example 32, the impregnated catalyst was used, while in Example 33 the impregnated catalyst was used. % By volume of sulfur dioxide

<td>Catalyst metal content by weight%</td><td>% Conversion of carbon monoxide</td>

<td> 1,0</td><td> 46</td>

<td> 2,0</td><td> 51</td>

<td> 0,3</td><td> 34</td>

<td> 1,0</td><td> 35</td>

<td> —</td><td> 28</td>

<td> 1</td><td> 65</td>

<td> ·—</td><td> 33</td>

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Examples 34 and 35

The procedure of Examples 32 and 33 in the Comparative Examples was repeated using mixtures of molecular sieve cracking catalyst and different amounts of alumina as prepared in Example 12 with impregnation with vanadium and copper. In Example 34, 10 g of vanadium and copper impregnated alumina catalyst are mixed with 90 g of a conventional cracked equilibrium catalyst containing 3.3% by weight molecular sieves, the catalyst is removed from a conventional catalytic moving bed cracker and calcined. 5 hours at 538 ° C. The synthetic waste gas flow rate (measured at 15.6 ° C) was 854 m³ per minute.

In Example 35, 0.5 g of vanadium and copper impregnated aluminum oxide of Example 12 was mixed with 49.5 g of the same calcined equilibrium and conventional cracking catalyst used in Example 34. The synthetic flue gas flow rate was 513 m.³ per minute (measured at 15.6 degrees Celsius) The volume percentages of sulfur dioxide removed from the waste gases of Examples 34 and 35 are shown in Table 2.

Examples 36 to 40

The procedure of Examples 32 and 33 was repeated using a regeneration temperature of 677 ° C. Example 36 is a comparative example using a synthetic smoke mixture flow rate of 1084 ml / min using the impregnated catalyst of Examples 13 and 14, while in Examples 37 and 38 the impregnated catalyst of Examples 13 and 14 was used and the synthetic flow rate of the of the exhaust gas mixture was 989 or 1 014 ml per minute, respectively. Example 39 is a comparative example using the non-impregnated catalyst of Examples 15, 34 and 35 at a flow rate of 891 ml of synthetic waste gas per minute. The procedure of Example 40 was carried out with the catalyst prepared in Example 15 at a flow rate of 992 ml of synthetic waste gas mixture per minute. All gas velocities were measured at 15.6 ° C. The percentages by volume of sulfur dioxide removed from the waste gases in Examples 36 to 40 are given in Table 2.

Examples 41 and 42

In Example 41, a 1.67 wt% sulfur oil feed gas cracked in a conventional moving bed cracker in an ascending reactor. The usual regeneration was performed, and a conventional equilibrium cracking catalyst of the molecular sieve type was used, containing 2.5% by weight molecular sieves and about 0.6% sodium. In Example 42, a second gas oil feed with a sulfur content of 1.68% by weight is cracked in the same apparatus using the same recovery system and the same cracking catalyst, but the catalyst is still impregnated with magnesium and zinc. Magnesium and zinc are deposited on the catalyst with magnesium sulphonic acid salt and zinc dialkyldithiophosphoric acid salt in minute concentrations as an additive to the lubricating oil in the feed to the reaction zone. After a few hours of this addition, the catalyst was found to contain 0.3 wt% magnesium and 0.1 wt% zinc. The operating conditions and the composition of the waste gases from the regeneration zone are given in Table 3.

Examples 43 and 44

Cracking of the gas oil with a sulfur content of 0.26% by weight is carried out in a conventional catalytic cracking apparatus using a moving bed in an ascending reactor and using conventional regeneration. In Example 44, a conventional equilibrium catalyst for moving bed cracking of molecular sieve type containing 2.5% by weight molecular sieve and 1.01% sodium was used. The sodium is deposited on the catalyst by means of an aqueous solution of sodium chloride added to the feed.

In Comparative Example 43, a gas oil with a sulfur content of 0.28% by weight was used in the presence of the same catalyst as in Example 44, but without sodium impregnation; otherwise, with identical identities. The operating conditions and composition of the gas mixture from the regeneration zone are given in Table 3. Examples 45 and 46

In Example 45, a 0.81 percent sulfur oil feed gas cracked in a conventional moving-bed catalytic cracking apparatus in an ascending reactor using the regeneration apparatus depicted in FIG. 2. The regeneration is carried out according to a regeneration scheme as shown in FIG. U.S. Patent 3,909,392. A conventional equilibrium cracking catalyst for the mobile phase cracking of molecular sieve type containing 4.5% by weight molecular sieve, 0.64% by weight iron, 56 ppm by weight copper and 0.22% by weight sodium was used. In Example 46, a second gas oil feed with a sulfur content of 1.14% by weight is used and is cracked in the same apparatus using the same regeneration system, and the equilibrium catalyst used for mobile phase cracking is a molecular sieve type, 2.5% molecular sieve, 0.52% iron, 34 ppm copper, 0.22% sodium and 0.09 ppm platinum. 272.1 tonnes of catalyst are used in the apparatus, and platinum is used

40 deposited on the catalyst with a solution of platinum acetylacetonate in benzene together with the reaction feed. A total of 20 g of platinum is used as a metal with a mean consumption of 3 g of platinum per day. In the example of both catalysts, sodium was already on the catalyst, iron and

Table 3

<td>Examples</td><td> 41</td><td> 42</td><td></td><td> 43</td><td> 44</td><td> 45</td><td> 46</td>

<td>Cracking conditions</td><td></td><td></td><td></td><td></td><td></td><td></td><td></td>

<td>Cracking temperature ° C</td><td> 524</td><td> 521</td><td></td><td> 493</td><td> 496</td><td> 509</td><td> 499</td>

<td>total feed rate</td><td></td><td></td><td></td><td></td><td></td><td></td><td></td>

<td>gas oil in hl / day</td><td> 5 866</td><td> 5 580</td><td></td><td> 2 760</td><td> 2 977</td><td> 6 443</td><td> 5 735</td>

<td>Volume ratio of total</td><td></td><td></td><td></td><td></td><td></td><td></td><td></td>

<td>spraying fresh</td><td> 1,07</td><td> 1,25</td><td></td><td> 1,12</td><td> 1,11</td><td> 1,07</td><td> 1,01</td>

<td>Speed ​​of circulation</td><td></td><td></td><td></td><td></td><td></td><td></td><td></td>

<td>of catalyst in kg / min</td><td> 26,4</td><td> 23,09</td><td></td><td> 16</td><td> 16</td><td> 22,6</td><td> 24,2</td>

<td>Weight ratio</td><td></td><td></td><td></td><td></td><td></td><td></td><td></td>

<td>catalyst to the feedstock</td><td> 5,3</td><td> 5,2</td><td></td><td> 7,2</td><td> 6,7</td><td> 4,2</td><td> 5,2</td>

<td>Conditions for ejection</td><td></td><td></td><td></td><td></td><td></td><td></td><td></td>

<td>Temperature, ° C</td><td> 521</td><td> 518</td><td></td><td> 491</td><td> 469</td><td> 512</td><td> 497</td>

<td>vapor to catalyst ratio</td><td></td><td></td><td></td><td></td><td></td><td></td><td></td>

<td>kg / kg</td><td> 4,3</td><td> 4,7</td><td></td><td> 4,6</td><td> 5,5</td><td> 4,1</td><td> 4,1</td>

<td>Regeneration conditions:</td><td></td><td></td><td></td><td></td><td></td><td></td><td></td>

<td>Thick bed temperature, ° C</td><td> 688</td><td> 707</td><td></td><td> 646</td><td> 619</td><td> 705</td><td> 719</td>

<td>Air speed</td><td></td><td></td><td></td><td></td><td></td><td></td><td></td>

<td>for combustion kg / h</td><td> 187 048</td><td> 221 098</td><td></td><td> 83 082</td><td> 81 265</td><td> 178 422</td><td> 206 570</td>

<td colspan="4">Composition of waste gases from the regeneration zone</td><td></td><td></td><td></td><td></td>

<td>CO2, mol. %</td><td> 13,0</td><td> 14,0</td><td></td><td> 11,8</td><td> 11,0</td><td> 17,4</td><td> 15,8</td>

<td>CO, mol. %</td><td> 5,0</td><td> 5,6</td><td></td><td> 7,3</td><td> 5,2</td><td> 0,1</td><td> 0,8</td>

<td>O2, mol. %</td><td> 0,5</td><td> 2,0</td><td></td><td> 0,5</td><td> 2,6</td><td> 1,0</td><td> 2,4</td>

<td>SO2, ppmv</td><td> 580</td><td> 60</td><td></td><td> 184</td><td> 34</td><td> 795</td><td> 627</td>

<td colspan="2">The results show that platinum is added</td><td>to</td><td colspan="2">oxide</td><td>aluminum,</td><td> 0,088 %</td><td>oxide</td>

<td>of the cracking catalyst</td><td>reaches lower</td><td>E-</td><td colspan="3">Silica, 0.005%</td><td>oxide</td><td>ferric</td>

<td colspan="2">mission of sulfur oxides, even if sprayed</td><td>at-</td><td>and</td><td> 0,004 %</td><td>oxide</td><td>sodium and</td><td>less than</td>

Hydrocarbons contain a greater sulfur content. In addition, the carbon monoxide emission was in each case less than 8 to 10 mol. %, which is a typical case for conventional regeneration in the absence of a catalyst with a platinum promoter. A certain increase in the carbon monoxide content of the off-gas from the regeneration zone using a cracking catalyst with a platinum promoter was in this case due to the higher velocity of the air to be burned. The temperature at the inlet to the first cyclone of Example 45 was 781 ° C and in Example 46 742 ° C. The drop in temperature in the inlet cyclone portion, the rise in bed temperature, and the reduction in the difference between the bed temperature and the inlet cyclone portions using a platinum cracking catalyst as promoter demonstrate a substantial increase in the combustion rate of carbon monoxide in the dense bed.

Example 47

Particulate alpha alumina monohydrate (CATAPAL-SB from Conoco Chemical Division of Continental Oil Company) containing 74.2% copper by weight is introduced into the cracking process cycle in the form of feed components. Working conditions and composition of waste gases from the regeneration zone are given in the table

3.

0.01% of sulfur having a density of 660 to 740 g / l, 78% of which have a particle diameter of less than 90 micrometers, is dried at 121 ° C and then calcined at 649 ° C for three hours, thereby the gamma aluminum oxide in the form of particles. Then, a solution of 56.4 g cadmium nitrate tetrahydrate in 240 ml water was used to impregnate 184.5 g of gamma alumina in particulate form. The impregnated alumina was dried at 121 ° C and then calcined at 649 ° C for three hours to give a particulate product containing 10% by weight of cadmium.

Example 48

The process described in Example 47 was repeated to impregnate 100 parts of gamma alumina in particulate form, but using a solution of 33.2 parts of a 51.8% manganese nitrate solution diluted with 112.8 parts of water. The impregnated alumina was dried and calcined as described in Example 47 to give a product containing 5 wt% manganese.

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Example 49

The process described in Example 47 is repeated, but a solution of ammonium paramiolybdate tetrahydrate of formula (ΝΗ4) βΜθ7θ24 is used to impregnate 25.1 g of particulate alumina. 4 H2O in 32.5 ml of water. The impregnated alumina was dried and calcined as described in Example 47 to give a product containing 10 wt% molybdenum. Example 50

The process described in Example 47 was repeated, but an aqueous solution of 36.1 parts of zinc nitrate hexahydrate was used to impregnate 100 parts of gamma alumina in particulate form. The impregnated alumina is dried and then calcined at 538 ° C for five hours to give a product containing 9% by weight of zinc, calculated as zinc oxide.

Example 51

The process described in Example 47 is repeated, but impregnating 100 parts of gamma alumina in particulate form is used.

47.5 parts of titanium sulphate of the formula

Ti 2 (SO 4) 3 dissolved in aqueous hydrogen peroxide solution. The impregnated alumina was dried, treated with hydrogen at 690 degrees Celsius and calcined at 538 degrees Celsius for 5 hours to give a particulate product containing titanium (9% by weight calculated as titanium dioxide).

Example 52

A solution of 10.8 g of sodium hydroxide in 100 ml of water is used to impregnate 81.3 g of alpha alumina in particulate form. The impregnated alumina was dried at 121 ° C and then calcined at 538 ° C for three hours to give the product as a sodium-containing particle (9% sodium by weight calculated as sodium oxide).

Example 53 g of amoma and cerium nitrate, containing

Dissolve 25.6% of cerium in 50 ml of water and use the resulting solution to impregnate 25.2 g of particulate silica gel. The impregnated silica gel is dried and then calcined at 538 ° C to give a particulate product containing 10 wt% cerium.

Example 54

Prepare a solution by dissolving 29.05 g of lanthanum nitrate hexahydrate, 9.94 g of lead nitrate and 35.80 g of a 50% aqueous manganese nitrate solution in sufficient water to give a total volume of 200 ml. The solution is then mixed with 60 ml of concentrated ammonium hydroxide solution. After standing overnight at room temperature, the precipitate formed is filtered off, washed with water, dried and calcined at 593.^CC. The resulting particulate product contains lanthanum, lead and manganese in an atomic ratio of 7: 3: 10.

Example 55

Prepare a solution by dissolving 3.80 g of copper nitrate trihydrate and 6.60 g of chromium nitrate nonahydrate in sufficient water to obtain 45 ml of the solution. This solution was used to impregnate 47.5 g of gamma alumina in the form of particles having a diameter of 833-1397 microns. The impregnated alumina is dried at 121 ° C for three hours and then calcined at 538 ° C for three hours to give the product as a copper-chromium-containing product, 2.5% by weight of copper oxide. and 2.5% chromium trioxide. The yellow, calcined product is ground to pass a 147 micron sieve.

Example 56

Prepare a solution by dissolving 0.3116 g of ferric nitrate nonahydrate in an amount of water to give 12 ml of the solution. The solution is then used to impregnate 20 g of alpha alumina in the form of particles passing through a sieve with holes of 147 microns and having a surface of 14²/G. The impregnated alumina is then dried at 121 ° C and then calcined at 538 ° C for three hours to give a particulate adduct containing 0.2% iron by weight.

Example 57

The process of Example 56 is repeated except that 20 g of alpha alumina is impregnated with a solution which is prepared by dissolving 0.3083 g of nickel nitrate hexahydrate in sufficient water to obtain 12 ml of solution. The impregnated alumina was dried and calcined under the conditions of Example 56 to give a particulate material containing 0.45% nickel by weight. Example 58

The process described in Example 24757 breaths 16-20 is repeated except that synthetic flue gas containing 1500 ppmv sulfur dioxide and 4% oxygen in nitrogen is passed through a solid fluidized catalyst bed and / or a particulate additive in the regenerator at temperature at either 677 or 788 ° C at a rate of 1,000 milliliters per minute. A 100 g sample of catalyst and / or kit is used for each test and the sulfur dioxide content of the discharged gas is measured with an ultraviolet analyzer. Before use, the ingredients are steamed for 5 hours at 760 ° C (100% steam at atmospheric pressure J.

The results for a CBZ-1 particulate cracking catalyst (Davison Chemical Division, WR Grace and Co.) having a content of 29.1% alumina, 0.46% sodium oxide and 0.11% iron and containing Y-zeolite are shown. in Table 4 as tests A and E.

The results for the various mixtures of CBZ-1 catalyst with 1 to 5% by weight of the particulate additive prepared as described in Example 47, 48 and 50 to 52 are shown in Table 4 as tests F, G and B to D. The results with the Filtrol 800 particulate cracking catalyst (molecular sieve catalyst containing 2% rare earth oxides and 46% alumina) by weight and the results of a Filtrol 800 mixture with 0.1% by weight particulate additive prepared as described in Example 54 are shown. in Table 4 as Test I and J. The results with the particulate additive prepared according to Example 53 are shown in Table 4 as Test H.

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Example 59

Pilot cyclic fluid catalytic cracking pilot tests are conducted with West Texas gas oil with a wide boiling range and a sulfur content of 1.47% by weight. Comparative tests are performed with CBZ-1 particulate cracking catalyst (Davison Chemical Divisdn, WR). Grace and Co.] and also using a mixture of 99 parts by weight of CBZ-1 catalyst 4S with 1 part by weight of 5% manganese on alumina prepared as described in Example 48 except that the curing is carried out at 538 ° C. . The addition of manganese on alumina reduces the emission of sulfur dioxide in the off-gas from the regenerator from 763 to 275 ppmv, which is a 64% reduction. Comparative values ​​are given in Table 5.

Table 5

Exam

<td></td><td>AND</td><td>(B)</td>

<td>Catalyst:</td><td>CBZ-1</td><td>CBZ-1 plus 1%</td>

<td>Cracking conditions</td><td></td><td></td>

<td>reactor temperature, ° C</td><td> 510</td><td> 510</td>

<td>speed, g / min</td><td> 9</td><td> 9</td>

<td>Stripping conditions</td><td></td><td></td>

<td>temperature, ° C</td><td> 495</td><td> 495</td>

<td>steam, g / h</td><td> 16</td><td> 16</td>

<td>nitrogen, m³/ h</td><td> 0,031</td><td> 0,031</td>

<td>Regeneration conditions</td><td></td><td></td>

<td>temperature, ° C</td><td> 649</td><td> 649</td>

<td>oxygen concentration</td><td></td><td></td>

<td>in the waste gas,</td><td></td><td></td>

<td>% by volume</td><td> 2,0</td><td> 2,0</td>

<td>concentration of sulfur oxides</td><td></td><td></td>

<td>in waste gas, ppmv</td><td> 763</td><td> 275</td>

The sum of sulfur dioxide and sulfur trioxide discharged from the regenerator is considered