RU2797245C2 - Method, reactor and system for catalytic cracking of liquid oil products - Google Patents

Abstract

FIELD: catalytic cracking.

SUBSTANCE: invention relates to a method for the catalytic cracking of liquid petroleum products, comprising the step of bringing the original liquid petroleum products into contact with a catalytic cracking catalyst for reaction in a reactor containing a fluidized bed with transport in a rarefied phase and with a fast fluidized bed connected in series, and in the fast fluidized bed an axial solid fractionε catalyst is measured at the top, middle, and bottom of the reaction zone, divided into equal parts in the axial direction, in each case are not less than about 0.1 and not more than about 0.2, providing an axial solids fractionε catalyst, which is adjusted within the range from about 0.1 to about 0.2 from bottom to top. The process is used to produce light olefins from heavy feedstocks and further comprises the steps of: i) contacting a first feedstock containing a light feedstock and/or a heavy feedstock with a fluid catalytic cracking catalyst to transport in a rarefied phase to carry out the first catalytic cracking reaction; and ii) introducing the reactor effluent from step (i) and optionally a second feed containing light feed and/or heavy feed oil into a second fast fluid catalytic cracking reaction under effective conditions to produce light olefins, wherein the axial solid factionε catalyst in a fast fluidized bed regulate within the range from about 0.1 to about 0.2 from the bottom up; iii) separating the effluent from the reactor to obtain a light olefin rich reaction product and a spent catalyst; iv) regenerating the spent catalyst and recycling at least a portion of the resulting regenerated catalyst to step (i) as a catalytic cracking catalyst; and v) optionally separating the reaction products to produce dry gas, liquefied petroleum gas, gasoline, diesel fuel and catalyst slurry. The invention also relates to a system suitable for use in the catalytic cracking of liquid petroleum products.

EFFECT: low yield of dry gas and coke and proper distribution of products.

17 cl, 2 dwg, 2 tbl, 4 ex

Description

Cross-reference to related applications

The present application claims the priority of PRC Patent Application No. 201810779819.0 entitled "Method and system for catalytic cracking using a fluidized bed with transport in a rarefied phase and a fast fluidized bed", filed on July 16, 2018, which is incorporated herein in its entirety by links.

TECHNICAL FIELD OF THE INVENTION

The present application relates to the field of catalytic cracking technology, in particular, to a method and system for catalytic cracking of liquid petroleum products.

State of the art of the present invention

Light olefins such as ethylene and propylene are the main raw materials for the chemical industry. In China and abroad, natural gas or light oil fraction is mainly used as a feedstock for the production of light olefins using the steam cracking process in an integrated ethylene production plant. Benzene, toluene and xylene (BTX) are important major chemical feedstocks, of which para-xylene (XX) accounts for approximately 45% of total BTX consumption. With the development of industries such as the polyester industry in the PRC, the demand for VTX is expected to continue to grow at a high rate. Approximately 90% ethylene, approximately 70% propylene, 90% butadiene and 30% aromatics provide steam cracking by-products. Although steam cracking technology has been developed for several decades and this technology is continuously improved, steam cracking technology still has the technical limitations of high energy consumption, high production cost, large CO ₂ emissions, difficult to control distribution of products, etc. When the conventional route for producing ethylene and propylene using steam cracking in the petrochemical industry is adopted, numerous limiting factors may occur, including shortage of light feed oil, insufficient production capacity, high cost, and so on. In addition, as steam cracking feedstocks become lighter, reduced yields of propylene and light aromatics increase the tension between supply and demand. Catalytic cracking technology can be used as a beneficial addition to the industrial process for producing light olefins and light aromatics and can provide clear social and economic benefits when used to produce chemical feedstocks in refineries and chemical process integrated plants.

PRC Patent Application Publication No. CN 1234426 A discloses a process for the simultaneous production of light olefins and high-aromatic gasoline from a heavy petroleum product, which involves introducing a heavy petroleum hydrocarbon and steam into a catalytic cracking reaction in a composite reactor, which is composed of a vertical reactor and a dense-phase fluidized layer, resulting in an increase in the yield of light olefins, in particular propylene, and the content of aromatic compounds in gasoline increases to approximately 80 wt. %.

PRC Patent Application Publication No. CN 1393510 A discloses a method for improving the production of ethylene and propylene by catalytically converting a heavy petroleum hydrocarbon, which involves contacting and reacting crude liquid petroleum products with a catalyst containing high-silica content pentasil-type zeolite in a vertical reactor or fluidized bed reactor. This method can not only increase the yield of ethylene and propylene, but also to some extent reduce the hydrothermal deactivation of the catalyst.

PRC Patent Application Publication No. CN 1721510 A discloses a process for producing light olefins and aromatics using catalytic cracking in two reaction zones, where different hourly volumes of oil per unit mass of catalyst are provided in the two reaction zones for the purpose of producing light olefins such as propylene. and ethylene, from heavy source oil products to the maximum extent, and the yield of propylene exceeds 20 wt. %, and the simultaneous production of aromatic compounds such as toluene and xylene.

US Patent Application Publication Nos. US 2002003103 A and US 2002189973 A disclose a fluid catalytic cracker (FCC) containing twin vertical reactors for improving propylene production, where gasoline (60-300°F/15-150°C), obtained in the cracking reaction enters the second vertical reactor for subsequent reaction, and the catalyst is a mixture of a USY type molecular sieve catalyst and a ZSM-5 type molecular sieve catalyst.

US Patent Application Publication No. US 2002195373 A and International Patent Application Publication No. WO2017223310A disclose processes using a downflow reactor operating at high temperature (1020-1200°F/550-650°C), short contact time ( less than 0.5 seconds) and a high ratio of catalyst and oil products (15-25). The primary catalyst (Faujasite Type Y) exhibits low hydrogen transfer activity and is designed to maximize light olefin yields in combination with operating conditions. A high efficiency separator is used to separate the product from the catalyst within 0.1 second, thus minimizing secondary reactions and coke formation. In addition, light cycle oil (LCO) is used to quench the separated product gas to a temperature of approximately 930°F/500°C and to prevent further cracking.

US Pat. No. 6,538,169 A and US Patent Application Publication No. US 2003121825 A also disclose methods using a reaction-regeneration system consisting of two reaction zones and a common regenerator. In the first reaction zone, the heavy feedstock is cracked to produce light olefins or intermediates that can be converted to light olefins at high temperature and high catalyst to oil ratio. The second reaction zone is a second vertical reactor where the process conditions are more stringent and the yield of light components from the gasoline feedstock is increased. A configuration selective molecular sieve such as ZSM-5 is used to facilitate the conversion of gasoline to light olefins and suitable feedstocks include vacuum gas oil (VGO), heavy vacuum gas oil (HVGO) and hydrogenated gas oil.

PRC Patent Application Publication No. CN 1403540 A discloses a catalytic conversion process for the production of ethylene and propylene, which uses a reactor composed of a vertical reactor and a dense phase fluidized bed connected in series, a light raw material is introduced into the vertical reactor to carry out the reaction under highly stringent conditions, and the reaction product and the coke-coated catalyst are sent to a fluidized bed for subsequent reaction under relatively mild conditions. This method can provide a relative increase in the total yield of ethylene, propylene and butene.

PRC Patent Application Publication No. CN 102051213 A discloses a catalytic cracking process which includes the steps of contacting a heavy feedstock with a catalyst in a first vertical reactor containing at least two reaction zones for the cracking reaction, and contacting a light feedstock and subjected to cracking heavy oil with a catalyst in the second vertical reactor and the reactor with a fluidized bed for carrying out the cracking reaction. When used for catalytic cracking of heavy oils, this process can provide a high degree of conversion of heavy oils and a high yield of propylene, as well as a low yield of dry gas and coke.

The structural contradiction of the petrochemical industry is becoming more and more serious. On the one hand, significant problems are the overproduction of traditional petrochemicals and the contradiction between the supply and demand of final petroleum products. On the other hand, there is also a significant shortage of high-grade petrochemical feedstocks. Thus, the transformation of oil refining into the chemical industry represents an irresistible trend. Catalytic crackers, which serve as bridges between the oil refining and chemical industries, must withstand unprecedented pressure and testing. At present, the proportion of atmospheric distillation residue in blended petroleum products that feed catalytic cracking units is constantly increasing, and there is even a proposal to introduce vacuum distillation residue into the blended petroleum products. Current catalytic cracking technology typically uses vacuum gas oil or atmospheric distillation residue based on paraffinic hydrocarbons as the feedstock, while the most advanced catalytic cracking technology uses a reactor that consists of twin vertical reactors or a vertical reactor connected in series with a dense phase bed in high rigidity conditions, for the purpose of increasing the production of light olefins and/or light aromatics. The problem of high dry gas and coke yields inevitably arises when such a reactor is used to process residue-containing heavy oil mixtures. A reduction in coke yield can be achieved by using a downflow reactor, but the conversion rate in the reaction is relatively low and a specialized catalyst is required. As the feed becomes heavier, the demand for mixed feeds containing residual oil products for catalytic cracking units increases. In order to efficiently use the resources of low grade heavy oil products and meet the growing demand for chemical feedstocks such as light olefins and heavy aromatics, it becomes necessary to develop a catalytic cracking process for converting a feedstock containing low grade heavy oil products into high value-added products.

Brief summary of the present invention

The object of the present application is to provide a new process and system for the catalytic cracking of liquid petroleum products which is particularly suitable for use in the production of light olefins such as ethylene and propylene from liquid petroleum feedstocks, in particular heavy petroleum feedstocks by means of catalytic cracking. . The catalytic cracking process and system has the advantage of low dry gas and coke yield and good product distribution.

In order to solve the problem presented above, according to one aspect of the present application, a process is provided for the catalytic cracking of liquid petroleum products, including the step of introducing the initial liquid petroleum products, in particular, heavy initial petroleum products, into contact with a catalytic cracking catalyst for reaction in a reactor containing a fluidized bed with transport in a rarefied phase and a fast fluidized bed, connected in series, and in the fast fluidized bed, the axial solid fraction s of the catalyst is controlled within the range from about 0.1 to about 0.2.

According to another aspect of the present application, a system is provided suitable for use in the catalytic cracking of liquid petroleum products, in particular heavy crude oils, and comprising a catalytic cracking reactor, a catalyst separating device, an optional reaction product separator, and a regenerator,

wherein the catalytic cracking reactor comprises a fluidized bed with a rarefied phase transport and a fast fluidized bed connected in series, wherein the fluidized bed with the rarefied phase transport is in fluid transfer communication with the fast fluidized bed and is located upstream of the relatively fast fluidized bed along the direction of movement of the reaction stream;

a fluidized bed with transport in a rarefied phase contains a catalyst inlet in the lower part and an inlet of the first source material in the lower part, a fast fluidized bed contains an outlet in the upper part and an optional inlet of the second source material in the lower part, the catalyst separating device contains an inlet, an outlet of the catalyst and an outlet reaction products, an optional reaction product separator comprises a reaction products inlet, a dry gas outlet, a liquefied petroleum gas (LPG) outlet, a gasoline outlet, a diesel fuel outlet, and a catalyst slurry outlet, and a regenerator comprises a catalyst inlet and a catalyst outlet,

wherein the inlet of the fluidized bed catalyst with the rarefied phase transport is in fluid transfer communication with the regenerator catalyst outlet, the fast fluidized bed outlet is in fluid transfer communication with the inlet of the catalyst separating device, the outlet of the reaction products of the catalyst separating device is in in fluid transfer communication with the reaction product inlet of the optional reaction product separator, and the catalyst outlet of the catalyst separating device is in fluid transfer communication with the regenerator catalyst inlet.

According to the present application, the density of the catalyst in the fast fluidized bed can be effectively increased by adjusting the axial solid fraction s of the catalyst in the fast fluidized bed within the range of about 0.1 to about 0.2, so that the ratio of catalyst to feed liquid oils is greatly improved. during the reaction, and a relatively long residence time of the hydrocarbon and the catalyst is obtained, whereby sufficient contact and reaction can be achieved between the catalyst and the feed liquid oils, in particular low grade heavy oils. As a result, the rate of the conversion reaction can be increased, the yield of light olefins and light aromatics can be increased, the required reaction temperature can be lowered, and at the same time, the production of dry gas and coke can be effectively reduced, so that product distribution and product quality are improved.

In addition, according to the present application, the ratio of catalyst and oil products can be controlled in a wide range by introducing additional catalyst into the fast fluidized bed, which provides more active sites for the cracking reaction. At the same time, by introducing an additional catalyst, the flexibility of controlling the reaction temperature can be increased, and it becomes possible to effectively control the temperature gradient and catalyst activity in the fast fluidized bed.

By applying the method and system of the present application, high value-added chemical raw materials can be obtained from cheap low-grade heavy oils in petrochemical plants to the maximum extent, an improved way of integrating oil refining and chemical technology in refineries can be provided, the problem of shortage of petrochemicals can be solved. raw materials, and the social and economic benefits of the petrochemical industry can be improved.

Other characteristics and advantages of the present application will be described in detail herein in the section "Detailed Disclosure of the Present Invention" below.

Brief description of the figures

The figures that form part of the present description are presented to aid understanding of the present application and should not be construed as limiting. The present application may be construed with reference to the figures in conjunction with the "Detailed Disclosure of the Present Invention" section below. Among the figures:

in fig. 1 is a schematic diagram of a preferred embodiment according to the present application; And

in fig. 2 is a schematic diagram of another preferred embodiment according to the present application.

Description of symbols

Detailed disclosure of the present invention

Hereinafter, the present application will be described in detail herein with reference to the respective specific embodiments and accompanying figures. It should be noted that specific embodiments of the present application are presented solely for the purpose of illustration, but are not intended to be limiting in any way.

Any specific numerical values, including the end points of the numerical range, which are described in the context of this application, are not limited to the corresponding exact values, but should be construed as further covering all values around the specified exact value, for example, all possible values within ± 5% from the exact value. In addition, with respect to any numerical range described herein, arbitrary combinations may be made between the endpoints of the range, between each endpoint and any particular value within the range, or between any two particular values within the range to produce one or more new numerical ranges, and these new numerical ranges should also be considered as specifically described in this application.

According to the present application, a dilute phase transport fluidized bed and a fast fluidized bed are two reactor reaction zones connected in series, and thus can also be referred to as a dilute phase transport fluidized bed reaction zone and a fast fluidized bed reaction zone. ", respectively.

According to the present application, the term "fluidized bed with rarefied phase transport" has a meaning that is well known to those skilled in the art, and in particular means a fluidized bed in which catalyst particles form a rarefied phase in suspension in a fluid and will be carried away by the fluid that exits the fluidized bed.

According to the present application, the term "fast fluidized reaction zone" has a meaning that is well known to those skilled in the art, and in particular means a reaction zone in which the catalyst particles are present in a fast fluidized state. The fast fluidized state is a bubble-free fluidized contact state of gaseous and solid phases, an important characteristic of which is that solid particles tend to move in clusters. When the catalyst is present in a fast fluidized state, the axial solids fraction s of the catalyst in the reaction zone typically ranges from about 0.05 to about 0.4. However, in conventional fast fluidized beds, the catalyst is typically distributed in a rarefied to dense state from top to bottom, for example, the axial solids fraction ε of the catalyst at the top may range from about 0.05 to about 0.1, and the axial the solid fraction ε of the catalyst at the bottom can range from about 0.3 to about 0.4.

According to the present application, in a fast fluidized bed, when the axial solids fraction ε of the catalyst is adjusted within the range from about 0.1 to about 0.2 from the bottom up (i.e., the axial solids fraction ε of the catalyst, measured in the upper, middle and lower parts of the reaction zone divided into equal parts in the axial direction, in each case is not less than about 0.1 and not more than about 0.2), the catalyst is distributed in a quasi-uniform fully dense phase manner throughout the fast fluidized bed. Accordingly, a fast fluidized reaction zone in which the catalyst is distributed in such a fully dense phase manner may be referred to by the term "fully dense phase reaction zone".

According to the present application, the term "mass ratio of water and oil" means the ratio of the total mass of steam introduced into the reactor, and the mass of the source material.

According to the present application, the terms "upstream" and "downstream" are used in relation to the direction of movement of the reaction stream. For example, when the reaction stream moves from bottom to top, the term "upstream" means a lower position, and the term "downstream" means a higher position.

Unless otherwise indicated, terms used herein have the meanings commonly understood by those skilled in the art; and if terms are defined herein and their definitions differ from the common understanding in the art, then the definitions provided herein shall prevail.

In the context of this application, other than items that are specified in a certain way, any item or items that are not mentioned should be considered the same as items known in the art, without any change. In addition, any of the embodiments that are described herein may be freely combined with one or more other embodiments that are described herein, and the technical solutions or ideas resulting from this should be considered as part of the original disclosure. or the original description of the present application, but should not be considered as a new subject that was not described or predicted in this document, unless it is clear to specialists in the art that such a combination is obviously unreasonable.

All patent and non-patent documents cited herein, including, but not limited to, tutorials and scientific articles, are incorporated herein by reference in their entirety.

According to a first aspect of the present application, there is provided a process for the catalytic cracking of liquid petroleum products, comprising the step of contacting feedstock liquid oils, in particular heavy feedstocks, with a catalytic cracking catalyst for reaction in a reactor containing a fluidized bed with transport in a rarefied phase and a fast fluidized bed, connected in series, and in the fast fluidized bed, the axial solid fraction ε of the catalyst is controlled within the range from about 0.1 to about 0.2.

According to a preferred embodiment of the present application, the process is used to produce light olefins such as ethylene and propylene from a heavy petroleum feedstock, and the process further comprises the steps of:

i) contacting the first feedstock, comprising the light feedstock and/or the heavy feedstock, with a fluid catalytic cracking catalyst in a rarefied phase transport to carry out a first catalytic cracking reaction; And

ii) introducing the reactor effluent from step (i) and optionally a second feed containing light feed and/or heavy feed oil into a second fast fluid catalytic cracking reaction,

wherein the light feedstock is selected from the group consisting of C4 hydrocarbon cut, C5-C6 light gasoline cut, and combinations thereof, and at least one of the first and second feedstocks contains a heavy oil feedstock;

while in the fast fluidized bed, the axial solid fraction ε of the catalyst is controlled within the range from about 0.1 to about 0.2.

According to the present application, the axial solids fraction e of the catalyst in the fast fluidized bed is adjusted within the range from about 0.1 to about 0.2 from bottom to top, so that distribution of the catalyst from rarefied to dense state in the fast fluidized bed from top to bottom is prevented, the actual ratio catalyst and oil products are maintained at a steady level in the fast fluidized bed from top to bottom, the yield of dry gas and coke can be reduced, and the yield of the desired product can be increased.

According to a further preferred embodiment of the present application, the method further comprises the following steps:

i) contacting the first feedstock with a rarefied phase fluidized catalytic cracking catalyst to carry out the first catalytic cracking reaction;

ii) introducing the reactor effluent from step (i) and optionally the second feedstock into a second fast fluid catalytic cracking reaction under effective conditions to produce light olefins, wherein the axial solids fraction s of the fast fluidized bed catalyst is controlled within a range of about 0.1 to about 0.2;

iii) separating the effluent from the reactor to obtain a light olefin rich reaction product and a spent catalyst;

iv) regenerating the spent catalyst and recycling at least a portion of the resulting regenerated catalyst to step (i) as a catalytic cracking catalyst; And

v) optionally separating the reaction products to produce dry gas, liquefied petroleum gas, gasoline, diesel fuel and catalyst slurry.

As used herein, the term "effective for the production of light olefins" means that at least a portion of the feedstock is subjected to efficient cracking, for example, deep catalytic cracking, in a fast fluidized bed to obtain light olefins, such as ethylene and propylene, as products, resulting in a product mixture enriched in light olefins.

As used herein, the expression "enriched in light olefins" means that the total content of light olefins (eg, ethylene and propylene) in the product or mixture of reaction products is more than about 10 wt. %, preferably more than about 15 wt. % and more preferably more than about 20 wt. % product mixture of reaction products.

In some embodiments, prior to said step (i) and/or after said step (ii), the process of the present application may further comprise one or more additional reaction steps such as catalytic cracking and/or catalytic isomerization carried out in an additional reaction zone with fluidized bed such as rarefied phase transport fluidized bed, dense phase fluidized bed, traditional fast fluidized bed, etc.

According to some other embodiments of the present application, the method does not include any additional reaction step before said step (i) or after said step (ii).

According to some particularly preferred embodiments of the present application, the method further comprises the following steps:

i) introducing a first feedstock containing a preheated low grade heavy oil product into the bottom of the vacuum fluidized bed to contact the catalytic cracking catalyst and introduce it into the first catalytic cracking reaction in the process of passing through the vacuum fluidized bed. a bottom-up phase to obtain an effluent from the reactor containing the product of the first reaction and the semi-spent catalyst;

ii) introducing the reactor effluent from step (i) into the bottom of the fast fluidized bed and introducing it into the second catalytic cracking reaction, together with an optional second feed containing light feedstock introduced from the bottom of the fast fluidized bed, while passing through the fast fluidized bed. layer from bottom to top to obtain an effluent from the reactor containing the product of the second reaction and the spent catalyst, and the axial solid fraction s of the catalyst in the fast fluidized bed is regulated so as to satisfy the condition: 0.1<ε<0.2;

iii) separating the product of the second reaction from the spent catalyst in the reactor effluent after step (ii);

iv) sending the spent catalyst to a regenerator for regeneration by coke burnout and recycling at least a portion of the resulting regenerated catalyst to step (i) as a catalytic cracking catalyst; And

v) optionally separating the resulting second reaction product to produce dry gas, liquefied petroleum gas, gasoline, diesel fuel and a catalyst slurry.

According to some other particularly preferred embodiments of the present application, the method further comprises the following steps:

i) introducing the first feedstock containing the light feedstock into the bottom of the rarefied phase transport fluidized bed to contact the catalytic cracking catalyst and introducing it into the first catalytic cracking reaction while passing through the fluidized bed with the rarefied phase transport from below up to obtain an effluent from the reactor containing the product of the first reaction and semi-spent catalyst;

ii) introducing the effluent from the reactor from step (i) into the bottom of the fast fluidized bed and introducing it into the second catalytic cracking reaction along with a second feed containing preheated low grade heavy oil introduced from the bottom of the fast fluidized bed while passing through the fast fluidized bed from bottom to top to obtain an effluent from the reactor containing the product of the second reaction and the spent catalyst, and the axial solid fraction ε of the catalyst in the fast fluidized bed is regulated so as to satisfy the condition: 0.1<ε<0.2;

iii) separating the product of the second reaction from the spent catalyst in the reactor effluent after step (ii);

iv) sending the spent catalyst to a regenerator for regeneration by coke burnout, and recycling at least a portion of the resulting regenerated catalyst to step (i) as a catalytic cracking catalyst; And

v) optionally separating the resulting second reaction product to produce dry gas, liquefied petroleum gas, gasoline, diesel fuel and a catalyst slurry.

According to a preferred embodiment of the present application, the process further comprises introducing one or more additional catalyst streams into the fast fluidized bed and contacting the additional catalysts with material in the fast fluidized bed to carry out the catalytic cracking reaction.

According to the present application, a single or each of several streams of additional catalyst may independently have a carbon content of approximately 0 to 1.0 wt. %. For example, the single or multiple additional catalyst streams may be independently selected from reclaimed catalysts, spent catalysts, and semi-reclaimed catalysts, including reclaimed, spent, and semi-reclaimed catalytic cracking catalysts.

According to the present application, the total amount of one or more streams of additional catalyst may be from approximately 0 to 50 wt. %, preferably from about 5 to 30 wt. % with respect to the catalyst circulation rate in the reactor.

According to the present application, a single or each of several streams of additional catalyst is independently introduced at a position at a distance from the bottom of the fast fluidized bed, constituting from about 0 to 90% of its full height. Preferably, the single or multiple additional catalyst streams are independently introduced at a level of about 20 to 80%, more preferably about 30 to 75% of the height of the fast fluidized bed. For example, the insertion position may be at the bottom of the fast fluidized bed or at about 1/3 of the full height of the fast fluidized bed. The temperature of the additional catalyst may be adjusted according to the desired reaction temperature, for example, cold and/or hot regenerated catalyst may be introduced, or cold and/or hot spent catalyst may be introduced.

According to the present application, by introducing additional catalyst into a fully dense phase reaction zone, the ratio of catalyst to oil can be controlled over a wider range and more active sites for the cracking reaction can be present. At the same time, by introducing an additional catalyst, the flexibility of controlling the reaction temperature can be increased, and it becomes possible to effectively control the temperature gradient and catalyst activity in the fully dense phase reaction zone. In addition, by introducing an additional catalyst into the fully dense-phase reaction zone, it is possible to maintain the highest possible density uniformity of the catalyst in the reaction zone, it becomes possible to effectively control the density distribution of the catalyst, so that the cracking reaction can be carried out completely and efficiently, and the selectivity of the desired product.

According to the present application, the distribution of the catalyst in the fast fluidized bed can be further controlled by adjusting the steam velocity in the fast fluidized bed and/or by using a distributor plate for the fast fluidized bed catalyst, whereby it becomes possible to distribute the catalyst in a quasi-homogeneous fully dense phase manner.

According to the present application, the original liquid oil products, such as a heavy oil product, in particular, a low-grade heavy oil product, can be introduced into a reactor containing a fluidized bed with transport in a rarefied phase and a fast fluidized bed, in one or more positions. For example, the original liquid oils can be introduced into a fluidized bed with transport in a rarefied phase entirely at one injection position or into a fast fluidized bed entirely at one injection position. Optionally, the feed liquid oils may be introduced into the rarefied phase transport fluidized bed and/or the fast fluidized bed at two or more injection positions in the same or different proportions. For example, a portion of the feed liquid oils may be introduced into the rarefied phase transport fluidized bed at one injection position, and another portion of the feed liquid oils may be introduced into the fast fluidized bed at a different injection position, or the feed liquid oils may be introduced into the fluidized bed. with transport in a rarefied phase at two or more injection positions, or the original liquid oils can be introduced into the fast fluidized bed at two or more injection positions.

According to the present application, a low grade heavy oil product means a heavy oil product that is less suitable for use in catalytic cracking processing than traditional heavy oil products. For example, the properties of a low grade heavy oil product may satisfy at least one, including one, two, three or four of the following criteria: density at 20°C is from about 900 to 1000 kg/m ³ , preferably from about 910 to 940 kg / m ³ ; the carbon residue is approximately from 2 to 10 wt. %, preferably from about 3 to 8 wt. %; the total nickel and vanadium content is from about 2 to 30 ppm, preferably from about 5 to 20 ppm; the characteristic factor K is less than about 12.1, preferably less than about 12.0. Carbon residue in a low grade heavy oil product is measured by the Conradson carbon residue test according to ASTM D-189.

For example, the low grade heavy oil may be heavy petroleum hydrocarbons and/or other mineral oils; heavy oil hydrocarbon may be selected from the group consisting of vacuum distillation residue (VR), low grade atmospheric distillation residue (AR), low grade hydrogenated residue, coking gas oil, deasphalted oil, vacuum gas oil having a high acid number crude oil having a high metal content crude oil and combinations thereof; wherein said other mineral oil may be selected from the group consisting of coal liquefaction oil, oil sand bitumen, shale oil, and combinations thereof.

The catalytic cracking catalyst used in the present invention is not particularly limited, and may be, for example, a variety of catalytic cracking catalysts known to those skilled in the art and suitable for producing light olefins from liquid petroleum feedstocks such as heavy feedstocks. According to a preferred embodiment, the catalytic cracking catalyst contains, based on the dry weight of the catalytic cracking catalyst, from about 1% to about 50%, preferably from about 5% to about 45%, more preferably from about 10% to about 40% by weight of zeolite, from about 5% to about 99%, preferably from about 10% to about 80%, more preferably from about 20% to about 70% by weight of inorganic oxide and from about 0% to about 70%, preferably from about 5% to about 60 %, preferably from about 10% to about 50% by weight of the clay; wherein the active zeolite may comprise a mesoporous zeolite and optionally a macroporous zeolite. Preferably, the mesoporous zeolite may be from about 0 to 50 wt. %, preferably from about 0 to 20 wt. % in relation to the total weight of the zeolite in terms of dry weight.

According to the present application, mesoporous and macroporous zeolites are zeolites that are well known in the art, i. the mesoporous zeolite has an average pore size of about 0.5 to 0.6 nm, and the macroporous zeolite has an average pore size of about 0.7 to 1.0 nm.

For example, the macroporous zeolite may be one or more zeolites selected from the group consisting of rare earth substituted Y (REY) zeolites, rare earth substituted H-Y (REHY) zeolites, ultrastable Y zeolites, and high silica zeolites of the Y type. Y obtained in various ways. The mesoporous zeolite may be selected from zeolites having an MFI structure such as ZSM zeolites and/or ZRP zeolites. Optionally, the mesoporous zeolite may be modified with a non-metallic element such as phosphorus and/or a transition metal element such as iron, cobalt, nickel. A more detailed description of ZRP zeolites can be found in US Pat. No. US5,232,675A. The ZSM zeolite is preferably a mixture of one or more zeolites selected from ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38, ZSM-48 and other zeolites having a similar structure. A more detailed description of ZSM-5 zeolites can be found in US Pat. No. US3,702,886A.

According to the present application, the inorganic oxide acting as a binder is preferably silicon dioxide (SiO ₂ ) and/or alumina (Al ₂ O ₃ ). The clay acting as the matrix (ie the carrier) is preferably kaolin and/or halloysite.

The catalytic cracking reaction conditions used according to the present invention are not strictly limited and may be, for example, conditions that are well known to those skilled in the art and suitable for producing light olefins from liquid petroleum feedstocks such as heavy feedstocks. According to a preferred embodiment, the first catalytic cracking reaction can be carried out under the following conditions: reaction temperature is from about 500 to 600°C, reaction time is from about 0.05 to 5 seconds, the weight ratio of catalyst to oil is from about 1:1 to about 50:1, water to oil weight ratio from about 0.03:1 to about 0.5:1, catalyst density from about 20 to 100 kg/m ³ , steam velocity from about 4 to 18 m/s , the reaction pressure is about 130 to 450 kPa, and the mass flow rate G _s of the catalyst flow is about 180 to 500 kg/(m ² s). More preferably, the first catalytic cracking reaction may be carried out under the following conditions: the reaction temperature is from about 520 to 580° C., the reaction time is from about 1 to 3 seconds, the mass ratio of catalyst to oil is from about 5:1 to about 25:1, the mass ratio of water and oil is from about 0.05:1 to about 0.3:1.

According to a preferred embodiment, the second catalytic cracking reaction can be carried out under the following conditions: reaction temperature is from about 510 to 650°C, reaction time is from about 1 to 20 seconds, the weight ratio of catalyst to oil is from about 3:1 to about 50 :1, water to oil weight ratio is from about 0.03:1 to about 0.8:1, catalyst density is from about 120 to 290 kg/m ³ , steam velocity is from about 0.8 to 2.5 m /s, the reaction pressure is about 130 to 450 kPa, and the mass flow rate Gs of the catalyst is about 15 to 150 kg/(m ² s). More preferably, the second catalytic cracking reaction can be carried out under the following conditions: the reaction temperature is from about 550 to 620° C., the reaction time is from about 3 to 15 seconds, the mass ratio of catalyst to oil is from about 10:1 to about 30:1, the mass ratio of water and oil is from about 0.05:1 to about 0.5:1, the density of the catalyst is from about 150 to 250 kg/m ³ , the vapor velocity is from about 1 to 1.8 m/s, the reaction pressure is about 130 to 450 kPa, and the mass flow rate G _s of the catalyst flow is about 20 to 130 kg/(m ² s).

According to the present application, the separation of the reaction product from the spent catalyst can be carried out in a manner known to those skilled in the art, for example, in a separator using a cyclone. The method of subsequent separation of the reaction products to obtain dry gas, liquefied petroleum gas, naphtha, diesel fuel and catalyst slurry is also well known to those skilled in the art. According to a preferred embodiment, the dry gas and LPG can then be separated using a separation device commonly used in the art to obtain the desired product such as ethylene, propylene, etc.

According to a preferred embodiment of the present application, the method further comprises introducing a C4 hydrocarbon fraction and/or a C5-C6 light gasoline fraction as a light feedstock comprising said first and/or second feedstocks into said fast fluidized bed and/or fluidized bed with rarefied phase transport at one or more positions for catalytic cracking reactions. For example, in some preferred embodiments, at least one of the first and second feedstocks comprises a light feedstock selected from the group consisting of C4 hydrocarbon cut, C5-C6 light gasoline cut, and combinations thereof. According to some additional preferred embodiments, the first feed comprises a light feed and a heavy feed oil, and at least a portion of the light feed is introduced into the fluidized bed with transport in rarefied phase upstream of the position at which the heavy feed oil is introduced into the fluidized bed. with transport in a rarefied phase. According to further preferred embodiments, the first feedstock comprises a heavy feedstock, such as a low grade heavy feedstock, and the second feedstock comprises a light feedstock. According to further preferred embodiments, the first feedstock comprises a light feedstock and the second feedstock comprises a heavy feedstock such as a low grade heavy feedstock.

According to the present application, the term "C4 hydrocarbon fraction" means a light hydrocarbon fraction that is present in gaseous form at normal temperature and pressure and contains C4 hydrocarbons as a main component, including alkanes, alkenes and alkynes containing 4 carbon atoms in the molecule. This fraction may contain gaseous hydrocarbon products (for example, liquefied petroleum gas) obtained by the process according to the present application and enriched in the C4 hydrocarbon fraction, or gaseous hydrocarbons produced by other plants and enriched in the C4 fraction, preferably the C4 hydrocarbon fraction, obtained by the process according to the present application. The C4 hydrocarbon fraction is preferably an olefin-rich C4 hydrocarbon fraction that may have a C4 olefin content of greater than about 50 wt. %, preferably more than about 60 wt. %, preferably more than about 70 wt. %.

According to the present application, the term "C5-C6 light gasoline fraction" means a component of gasoline containing 5 or 6 carbon atoms, which may contain gasoline obtained in the cracking reaction of the method according to this application, and may also contain gasoline fractions produced by other installations, for example, at least one C5-C6 fraction selected from the group consisting of deep catalytic cracking gasoline, catalytic cracking gasoline, straight run gasoline, coking gasoline, thermal cracking gasoline and hydrogenated gasoline.

According to the present application, the spent catalyst can be regenerated using means well known to those skilled in the art, such as being regenerated in a regenerator by burning coke. In particular, an oxygen-containing gas such as air may be introduced into the regenerator to come into contact with the spent catalyst. The off-gas resulting from coke-burning regeneration can be separated from the catalyst in the regenerator and then sent to a downstream power recovery system.

According to some preferred embodiments of the present application, the regenerated catalyst resulting from regeneration by burning coke in the regenerator may be cooled to a temperature of about 600 to 680° C. in a catalyst cooler before being recycled to the reactor. Recirculation to the reactor after cooling the hot regenerated catalyst is advantageous for lowering the contact temperature of the oil and catalyst, improving the contact state of the feed oil and the catalyst, and further reducing the dry gas and coke selectivity.

According to some preferred embodiments, the rapid fluidized bed is comprised of bottom-up fully dense phase reaction zone and transition section, wherein the fully dense phase reaction zone is in the form of a hollow column having an approximately circular cross-section of constant diameter or variable diameter, an open lower end, and an open upper end. , a fluidized bed with rarefied phase transport is in communication with the lower end of the fully dense phase reaction zone, the upper end of the fully dense phase reaction zone is in communication with the outlet section of the reactor through a transition section, the lower section of the fully dense phase reaction zone optionally contains one or more inlets for introducing second source material,

wherein the cross-sectional diameter of the lower end of the fully dense-phase reaction zone is greater than or the same as the diameter of the fluidized bed with transport in rarefied phase and the cross-sectional diameter of the upper end of the fully dense-phase reaction zone is greater than the diameter of the outlet section, and the bottom or side wall a fully dense phase reaction zone contains one or more inlets for additional catalyst, each independently located at a level of from about 0 to 90%, preferably from about 20 to 80%, more preferably from about 30 to 75% relative to the total height fast fluidized bed.

According to the present application, a fully dense-phase reaction zone may be present in the form of a hollow column having an approximately circular cross-section with a constant diameter or a variable diameter, an open lower end and an open upper end, such as a hollow cylinder with a constant diameter or a hollow column whose diameter is continuous or incrementally increases from bottom to top.

According to the present application, the expression "diameter continuously increases" means that the diameter increases continuously in a linear or non-linear manner. An inverted hollow truncated cone can be mentioned as an example of a "hollow column whose diameter continuously increases from bottom to top".

According to the present application, the expression "diameter discontinuously increases" means that the diameter increases in a discontinuous manner, for example, increases in a stepped manner. As an example of a "hollow column whose diameter increases in a discontinuous manner from bottom to top", there may be mentioned a hollow column that consists of two or more cylindrical sections with successively increasing diameters.

For example, a fully dense-phase reaction zone may be present in the form of a hollow cylinder, an inverted hollow truncated cone, a hollow column consisting of two or more cylindrical sections with successively increasing diameters, a hollow column consisting of two or more sections of inverted truncated cones with successively increasing diameters, or a hollow column consisting of one or more cylindrical sections and one or more sections of inverted truncated cones.

In some preferred embodiments, the bottom section of the fully dense phase reaction zone contains a catalyst distributor plate.

In some preferred embodiments, the lower section of the fully dense phase reaction zone comprises one or more second feed inlets, and preferably a gas distributor is present at the second feed inlet position.

In a preferred embodiment, the ratio of the diameter of the maximum cross section of the fully dense phase reaction zone to the total height of the rapid reaction bed is from about 0.005:1 to about 1:1, preferably from about 0.01:1 to about 0.8:1, more preferably from about 0.05:1 to about 0.5:1; the ratio of the height of the fully dense phase reaction zone to the full height of the fast reaction bed is from about 0.1:1 to about 0.9:1, preferably from about 0.3:1 to about 0.85:1, more preferably from about 0.5 :1 to about 0.8:1.

In some preferred embodiments, the fully dense phase reaction zone is in the form of an inverted hollow truncated cone with an isosceles trapezoid longitudinal section that has a lower cross-sectional diameter of about 0.2 to 10 meters, preferably about 0.5 to 8 meters. , preferably about 1 to 5 meters; the ratio of the diameter of the upper cross section and the diameter of the lower cross section is from more than 1 to about 50, preferably from about 1.2 to about 10, more preferably from about 1.5 to about 5; the ratio of the maximum cross-sectional diameter to the total height of the fast reaction bed is from about 0.005:1 to about 1:1, preferably from about 0.01:1 to about 0.8:1, more preferably from about 0.05:1 to about 0 .5:1; the ratio of the height of the fully dense phase reaction zone to the full height of the fast reaction bed is from about 0.1:1 to about 0.9:1, preferably from about 0.3:1 to about 0.85:1, more preferably from about 0.5 :1 to about 0.8:1.

According to some other preferred embodiments, the fully dense-phase reaction zone is present in the form of a hollow column, which consists of an inverted truncated cone section and a cylindrical section, with the truncated cone preferably located below the cylinder, while the truncated cone has a longitudinal section in the form of an isosceles trapezoid, the diameter of the lower the cross section is about 0.2 to 10 meters, preferably about 0.5 to 8 meters, more preferably about 1 to 5 meters; and the ratio of the diameter of the upper cross section and the diameter of the lower cross section is from more than 1 to about 50, preferably from about 1.2 to about 10, more preferably from about 1.5 to about 5; the diameter of the cylinder is approximately the same as the diameter of the upper cross-section of the frustocone, and the ratio of the height of the hollow cylinder to the height of the frustum is from about 0.4:1 to about 2.5:1, preferably from about 0.8:1 to about 1.5:1; the ratio of the diameter of the maximum cross section of the fully dense phase reaction zone to the total height of the fast reaction bed is from about 0.005:1 to about 1:1, preferably from about 0.01:1 to about 0.8:1, more preferably from about 0.05: 1 to about 0.5:1; the ratio of the height of the fully dense phase reaction zone to the full height of the fast reaction bed is from about 0.1:1 to about 0.9:1, preferably from about 0.3:1 to about 0.85:1, more preferably from about 0.5 :1 to about 0.8:1.

In other preferred embodiments, the fully dense phase reaction zone is in the form of a hollow cylinder having a diameter of about 0.2 to 10 meters, preferably about 1 to 5 meters; the ratio of the diameter of the fully dense phase reaction zone to the total height of the fast reaction bed is from about 0.005:1 to about 1:1, preferably from about 0.01:1 to about 0.8:1, more preferably from about 0.05:1 to about 0.5:1; the ratio of the height of the fully dense phase reaction zone to the full height of the fast reaction bed is from about 0.1:1 to about 0.9:1, preferably from about 0.3:1 to about 0.85:1, more preferably from about 0.5 :1 to about 0.8:1.

In a preferred embodiment, the fully dense phase reaction zone has a height of about 2 to 50 meters, preferably about 5 to 40 meters, and more preferably about 8 to 20 meters.

In a preferred embodiment, the ratio of the height of the transition section to the overall height of the fast reaction bed is from about 0.1:1 to about 0.9:1, preferably from about 0.2:1 to about 0.5:1. Preferably, the transition section is in the form of a hollow truncated cone with a longitudinal section in the form of an isosceles trapezoid with sides having an internal inclination angle α of about 25 to 85°, preferably about 30 to 75°.

According to some embodiments of the present application, the reactor used may further comprise one or more additional fluidized beds, which are reaction zones and located upstream of a fluidized bed with transport in a rarefied phase and/or downstream of a relatively fast fluidized bed, such as a fluidized bed with rarefied phase transport, dense phase fluidized bed, traditional fast fluidized bed, etc.

According to some other embodiments of the present application, the reactor used does not have any additional reaction zone upstream of the rarefied phase transport fluidized bed or downstream of the relatively fast fluidized bed.

According to some variants of implementation of the present application, the reactor used may be located coaxially with the separator, or it may be located in parallel with the separator.

According to a second aspect of the present application, there is provided a system suitable for catalytic cracking of liquid petroleum products, in particular heavy crude oils, and comprising a catalytic cracking reactor, a catalyst separating device, an optional reaction product separator, and a regenerator,

wherein the catalytic cracking reactor comprises a fluidized bed with a rarefied phase transport and a fast fluidized bed connected in series, wherein the fluidized bed with the rarefied phase transport is in fluid transfer communication with the fast fluidized bed and is located upstream of the relatively fast fluidized bed along the direction of movement of the reaction stream;

a fluidized bed with transport in a rarefied phase contains a catalyst inlet in the lower part and an inlet of the first source material in the lower part, a fast fluidized bed contains an outlet in the upper part and an optional inlet of the second source material in the lower part, the catalyst separating device contains an inlet, an outlet of the catalyst and an outlet reaction products, the optional reaction product separator comprises a reaction products inlet, a dry gas outlet, an LPG outlet, a gasoline outlet, a diesel fuel outlet, and a catalyst slurry outlet, and the regenerator comprises a catalyst inlet and a catalyst outlet,

the fluidized bed catalyst inlet to the rarefied phase transport is in fluid transfer communication with the regenerator catalyst outlet, the fast fluidized bed outlet is in fluid transfer communication with the catalyst separating device inlet, the reaction products outlet of the catalyst separating device is in communication with a fluid transfer capability with a reaction product inlet of an optional reaction product separator, and a catalyst outlet of the catalyst separating device is in fluid transfer communication with a catalyst inlet of the regenerator.

According to a preferred embodiment, the fast fluidized bed and the rarefied phase transport fluidized bed are arranged coaxially one above the other, and the fast fluidized bed is located above the rarefied phase transport fluidized bed.

In some preferred embodiments, the fast fluidized bed comprises a catalyst distributor plate that can be located at the bottom of the fast fluidized bed, for example, at the junction of the rarefied phase transport fluidized bed and the fast fluidized bed.

According to the present application, the catalyst distributor plate may be any type of distributor plate that is commonly used in the industry, such as one or more flat, arcuate, cupped, annular and umbrella shaped plates. The use of a catalyst distributor plate is advantageous because it allows the catalyst to come into contact with the feed oil at a uniform concentration along the axial direction of the fully dense-phase reaction zone for the catalytic cracking reaction, thereby reducing the formation of vented coke and coke from thermal reaction due to excessively high or excessive low catalyst concentration.

In some preferred embodiments, the fast fluidized bed comprises one or more second feed inlets at the bottom, preferably with a gas distributor in the inlet position.

According to some preferred embodiments, the rapid fluidized bed has the structure as described above, i. a structure containing a completely dense-phase reaction zone and a transition section arranged sequentially from bottom to top. A detailed description of the specific configuration of the fully dense phase reaction zone and transition section will not be repeated here.

In some embodiments, one or more additional fluidized beds may be present in the catalytic cracking reactor, representing reaction zones, such as a fluidized bed with transport in a rarefied phase, a dense phase fluidized bed, a traditional fast fluidized bed, etc., located upstream flow relative to the fluidized bed with transport in a rarefied phase and/or downstream of the relatively fast fluidized bed.

In some other embodiments, the catalytic cracking reactor lacks any additional reaction zone located upstream of the rarefied phase fluidized bed or downstream of the relatively fast fluidized bed.

According to the present application, the catalyst separating device and the reaction product separator can be any device known to those skilled in the art. For example, the catalyst separating device may include a cyclone, separator, stripper, etc., and the reaction product separator may be a fractionator, etc.

The present application will be further described with the presentation of preferred embodiments, which are illustrated in the figures, but the present application is not limited to them.

In FIG. 1 shows a preferred embodiment according to the present application, in which the pre-lift medium, which may be a dry gas, steam or a mixture thereof, enters the bottom of the fluidized bed with transport in rarefied phase I through the bottom of the pre-lift section 2 through the pre-lift conduit medium 1. The regenerated catalyst from the regenerated catalyst riser 11 is introduced into the lower part of the pre-lifting section 2, then into the fluidized bed with transport in rarefied phase I, and moves up by the lifting action of the pre-lifting medium. The first feed containing C4 hydrocarbon cut, C5-C6 light naphtha and/or feed liquid oils such as low grade heavy oil enters the bottom of the fluidized bed with transport in rarefied phase I through the first feed line 14, mixes and enters the contacting a fluidized bed catalyst with rarefied phase transport I for the first catalytic cracking reaction to form an effluent from the reactor containing the product of the first reaction and the semi-spent catalyst. The effluent from the reactor moves upward to the bottom of the fast fluidized bed II, comes into contact with additional catalyst, which may be regenerated catalyst or spent catalyst, introduced through auxiliary conduit 15 for the second catalytic cracking reaction. Auxiliary conduit 15 is connected to the fast fluidized bed II at a level of about 0 to 90%, preferably about 20 to 80%, more preferably about 30 to 75% of the respective height. Optionally, a second feed containing C4 hydrocarbon cut, C5-C6 light naphtha and/or feed liquid oils such as low grade heavy oil enters the bottom of the fast fluidized bed II through the second feed conduit 16. The reactor effluent containing the product the second reaction and deactivated spent catalyst resulting from the reaction enters the cyclone 6 in the separator 4 through the outlet section 3 in order to separate the spent catalyst from the product of the second reaction. The separated product of the second reaction enters the injection chamber 7, and the reaction product in the injection chamber 7 enters the subsequent product separation system (not shown) through the reactor steam line 8. The fine particles of the catalyst are returned to the separator 4 through the dip riser of the cyclone 6, and the spent catalyst in the separator 4 enters the stripper 5. The reaction product, purified from the spent catalyst, enters the injection chamber 7 after passing through the cyclone 6. After the stripper, the spent catalyst enters the regenerator 10 through the vertical pipeline 9 of the spent catalyst, and the air enters the regenerator 10 after distribution through the air distributor 13 in for the purpose of burning coke on the spent catalyst in a dense bed at the bottom of the regenerator 10 such that the deactivated spent catalyst is regenerated to form a regenerated catalyst. The regenerated catalyst is recycled to the pre-lift section 2 via the regenerated catalyst riser 11, and the exhaust gas enters the downstream power recovery system (not shown) via the exhaust gas pipeline 12.

In FIG. 2 shows another preferred embodiment according to the present application, which is substantially identical to the embodiment shown in figure 1, except that the fast fluidized bed II consists of bottom up fully dense phase reaction zone 17 and transition section 18. Full dense phase reaction zone 17 is present in the form of an inverted hollow truncated cone with a longitudinal section in the form of an isosceles trapezoid. The transition section 18 is in the form of a hollow truncated cone with a longitudinal section in the form of an isosceles trapezoid with sides having an internal inclination angle α of about 25 to 85°, preferably about 30 to 75°.

According to some preferred embodiments, the following technical solutions are proposed in this application:

1. A process for catalytic cracking using a fluidized bed with transport in a rarefied phase and a fast fluidized bed, comprising the following stages:

i) introducing a preheated low grade heavy oil into the bottom of the fluidized bed with rarefied phase transport to contact the catalytic cracking catalyst and carry out the first catalytic cracking reaction in a bottom-up process to obtain a first reaction product and a semi-spent catalyst;

ii) introducing the resultant first reaction product and the semi-spent catalyst into the bottom of the fast fluidized bed to conduct the second catalytic cracking reaction in a bottom-up process to produce a second reaction product and a spent catalyst; moreover, the catalyst in the fast fluidized bed is distributed in a completely dense-phase manner, and the distribution of the axial solid fraction ε in the fast fluidized bed satisfies the condition: 0.1<ε<0.2;

iii) sending the spent catalyst to a regenerator for regeneration by coke burnout and recycling at least a portion of the resulting regenerated catalyst to the bottom of the fluidized bed with rarefied phase transport as a catalytic cracking catalyst; And

iv) separating the product of the second reaction to obtain dry gas, liquefied petroleum gas, gasoline, diesel fuel and catalyst slurry.

2. The method according to claim 1, in which the properties of low-grade heavy oil satisfy one, two, three or four of the following criteria: density at 20°C is from 900 to 1000 kg/m ³ , carbon residue is from 2 to 10 wt. %, the total content of nickel and vanadium is from 2 to 30 ppm, and the value of the characteristic factor K is less than 12.1.

3. The method according to claim 1, in which the properties of low-grade heavy oil satisfy one, two, three or four of the following criteria: density at 20°C is from 910 to 940 kg/m ³ , carbon residue is from 3 to 8 wt. %, the total content of nickel and vanadium is from 5 to 20 ppm, and the value of the characteristic factor K is less than 12.0.

4. The method according to paragraph 1, in which the low-grade heavy oil product is a heavy petroleum hydrocarbons and/or other mineral oils;

wherein the heavy petroleum hydrocarbon is one or more selected from vacuum distillation residue, low grade atmospheric distillation residue, low grade hydrogenated residue, coker gas oil, deasphalted oil, vacuum gas oil, high acid number crude oil, and high metal content. the crude oil, and wherein said other mineral oil is one or more of coal liquefaction oil, oil sand bitumen and shale oil.

5. The method according to paragraph 1, in which the catalytic cracking catalyst contains, in relation to the dry weight of the catalytic cracking catalyst, from 1 to 50 wt. % zeolite, from 5 to 99 wt. % inorganic oxide, and from 0 to 70 wt. % clay;

the zeolite contains a mesoporous zeolite, which is a ZSM zeolite and/or a ZRP zeolite, and optionally a macroporous zeolite, which is one or more materials selected as rare earth-substituted Y-type zeolite, rare-earth-substituted H-Y type zeolite, ultrastable type zeolite Y and high silica zeolite type Y.

6. The method according to paragraph 5, in which the mesoporous zeolite is from 0 to 50 wt. % relative to the total weight of the zeolite in terms of dry weight.

7. The method according to paragraph 5, in which the mesoporous zeolite is from 0 to 20 wt. % relative to the total weight of the zeolite in terms of dry weight.

8. The method according to claim 1, wherein the first catalytic cracking reaction is carried out under the following conditions: the reaction temperature is from 500 to 600°C, the reaction time is from 0.05 to 5 seconds, the mass ratio of catalyst and oil is from 1:1 to 50:1, the mass ratio of water and oil is from 0.03:1 to 0.5:1, the density of the catalyst is from 20 to 100 kg/m ³ , the steam velocity is from 4 to 18 m/s, the reaction pressure is from 130 to 450 kPa, and the mass flow rate Gs of the catalyst is from 180 to 500 kg/(m ² s);

the second catalytic cracking reaction is carried out under the following conditions, the reaction temperature is from 510 to 650°C, the reaction time is from 1 to 20 seconds, the mass ratio of catalyst and oil is from 3:1 to 50:1, the mass ratio of water and oil is from 0 03:1 to 0.8:1, the catalyst density is 120 to 290 kg/m3, the vapor velocity is 0.8 to 2.5 m/s, the reaction pressure is 130 to 450 kPa, and the mass velocity G _s of the catalyst flow is from 15 to 150 kg/(m ² s).

9. The method according to claim 1, wherein the first catalytic cracking reaction is carried out under the following conditions: the reaction temperature is from 520 to 580°C, the reaction time is from 1 to 3 seconds, the mass ratio of catalyst and oil is from 5:1 to 25: 1, the mass ratio of water and oil is from 0.05:1 to 0.3:1;

the second catalytic cracking reaction is carried out under the following conditions, the reaction temperature is from 550 to 620°C, the reaction time is from 3 to 15 seconds, the mass ratio of catalyst and oil is from 10:1 to 30:1, the mass ratio of water and oil is from 0 05:1 to 0.5:1, the density of the catalyst is 150 to 250 kg/m3, the vapor velocity is 1 to 1.8 m/s, and the mass flow rate G _{s of} the catalyst is 20 to 130 kg/( m ² s).

10. The process of claim 1, further comprising introducing a C4 hydrocarbon cut and/or a C5-C6 light gasoline cut into a fast fluidized bed and/or a rarefied phase transport fluidized bed for the catalytic cracking reaction.

11. The method of claim 10, wherein the C4 hydrocarbon cut and/or the C5-C6 light gasoline cut is introduced before the position where the low grade heavy oil product is introduced into the fluidized bed with transport in rarefied phase.

12. The method of claim 1, further comprising introducing additional catalyst into the fast fluidized bed to carry out the second catalytic cracking reaction together with the product of the first reaction and the semi-spent catalyst; while in which the carbon content of the additional catalyst is from 0 to 1.0 wt. %.

13. The method according to paragraph 12, in which the additional catalyst is from 0 to 50 wt. % of the total circulation rate of the fluidized bed catalyst with transport in a rarefied phase and fast fluidized bed.

14. The method according to paragraph 12, in which the additional catalyst is from 5 to 30 wt. % of the total circulation rate of the fluidized bed catalyst with transport in a rarefied phase and fast fluidized bed.

15. The method of claim 12, wherein the additional catalyst is introduced into the bottom of the fast fluidized bed.

16. A catalytic cracking system comprising a fluidized bed with transport in a rarefied phase, a fast fluidized bed, a catalyst separating device, a reaction product separator and a regenerator;

wherein the rarefied phase transport fluidized bed is in fluid transfer communication with the fast fluidized bed and is located upstream of the relatively fast fluidized bed along the direction of reaction flow;

fluidized bed with transport in rarefied phase contains a catalyst inlet in the lower part and an inlet of low-grade heavy oil product in the lower part, a fast fluidized bed contains an outlet in the upper part, a catalyst separating device contains an inlet, a catalyst outlet and an outlet of reaction products, a reaction product separator contains an inlet of products reactions, a dry gas outlet, an LPG outlet, a gasoline outlet, a diesel fuel outlet, and a catalyst slurry outlet, and the regenerator comprises a catalyst inlet and a catalyst outlet;

the fluidized bed catalyst inlet to the rarefied phase transport is in fluid transfer communication with the regenerator catalyst outlet, the fast fluidized bed outlet is in fluid transfer communication with the catalyst separating device inlet, the reaction products outlet of the catalyst separating device is in communication with a fluid transfer capability with the reaction product inlet of the reaction product separator, and the catalyst outlet of the catalyst separating device is in fluid transfer communication with the catalyst inlet of the regenerator.

17. The system of claim 16, wherein the fast fluidized bed is coaxial with the rarefied phase transport fluidized bed, and the fast fluidized bed is positioned above the rarefied phase transport fluidized bed.

The definition and calculation of each parameter in the present application is carried out as follows:

(1) Axial catalyst solids ε=pressure difference between two positions in the reaction zone along the axial direction, measured by a pressure difference sensor/distance between two positions along the axial direction/density of catalyst particles;

where the pressure drop is expressed in kg/m ² , the distance between two positions along the axial direction is expressed in meters (m), and the density of the catalyst particles is expressed in kg/m ³ .

Catalyst Particle Density=Theoretical Density / (Catalyst Pore Volume × Theoretical Density+1),

where the theoretical density is expressed in kg/m ³ and the pore volume of the catalyst is expressed in m ³ and the theoretical density and pore volume of the catalyst are determined by the methods of pycnometric study and water titration, respectively.

(2) Reaction time=reaction zone volume/average log oil vapor volume flow rate;

where the volume of the reaction zone is expressed in m ³ and the average logarithmic volumetric flow rate of the vapor of the oil product is expressed in m ³ /s.

Average logarithmic volume flow rate of oil vapor = (V _out -V _in ) / ln(V _out /V _in ),

where V _out and V _in are the volumetric flow rates of vapor oil at the outlet and inlet of the reaction zone, respectively;

the volumetric flow rate of steam oil at the outlet of the reaction zone V _out =m/ρ ₃ ;

the volumetric flow rate of steam oil at the inlet of the reaction zone V _in =m/ρ ₄ ;

where m is the input mass of the original oil and atomizing steam per unit time, expressed in kg/s; ρ ₃ is the oil vapor density at the outlet of the reaction zone, expressed in kg/m ³ ; ρ ₄ is the oil vapor density at the inlet of the reaction zone, expressed in kg/m ³ .

(3) Catalyst density in the reaction zone (or its top, middle, or bottom)=pressure difference between two positions in the reaction zone (or its top, middle, or bottom) along the axial direction, as measured by a differential pressure transducer/distance between the two positions along the axial direction;

where the pressure drop is expressed in kg/m ² , the reaction zone is divided into three equal parts, representing the upper part, the middle part and the lower part, along the axial direction, and the distance between the two positions along the axial direction is expressed in meters (m).

(4) Vapor velocity=average log volumetric oil vapor flow rate/cross-sectional area of the reaction zone;

when the reaction zone is not cylindrical, the vapor velocity means the average logarithmic value of the steam velocity at the bottom of the reaction zone and the steam velocity at the top of the reaction zone.

(5) Catalyst mass flow rate G _s = reactor catalyst circulation rate / cross-sectional area of the reaction zone;

when the reaction zone has a non-cylindrical shape, the mass flow rate G _{s of} the catalyst flow means the average logarithmic value of G _s at the bottom of the reaction zone and G _s at the top of the reaction zone;

where the catalyst circulation rate is expressed in kg/s;

reactor catalyst circulation rate=coke formation rate/(carbon content of spent catalyst - carbon content of regenerated catalyst);

where the rate of formation of coke is expressed in kg/s, and the carbon content of the spent catalyst and the carbon content of the regenerated catalyst are mass content;

coke formation rate=amount of exhaust gas/(CO ₂ %+CO%)/V _m ×M;

where V _m is the molar volume of the gas, which is 22.4×10 ^-3 m ³ /mol, and M is the molar mass of the element carbon, which is 12×10 ^-3 kg/mol;

exhaust gas quantity=(air velocity × 79 vol%)/(1 - CO ₂ % - CO % - O ₂ %);

where the air velocity is expressed in m ³ /s, the amount of exhaust gas is expressed in m ³ /s, and CO ₂ %, CO% and O ₂ % represent the volume percentage of CO ₂ , CO and O ₂ in the exhaust gas, respectively.

Examples

The present application will be further illustrated by presenting the following examples, but the present application is not limited to them.

All starting oils used in the following examples and comparative examples are hydrogenated residue, the properties of which are shown in Table 1. The catalyst used is an industrial catalytic cracking catalyst, which was purchased under the trade name DMMC-2 from a catalyst subsidiary of Sinopec Co., Ltd. .

Example 1

Example 1

This example was carried out in accordance with the process flow diagram shown in FIG. 1 in a medium size plant using a hydrogenated residue as the feed oil, a DMMC-2 catalyst as a catalyst, a combined reactor containing a rarefied phase transport fluidized bed and a fast fluidized bed as a reactor in series. The preheated feed oil was introduced into the fluidized bed with rarefied phase transport to contact the catalyst for the first cracking reaction. The effluent from the reactor moved upward into the fast fluidized bed to mix with additional regenerated catalyst for the second catalytic cracking reaction. The carbon content of the additional regenerated catalyst was 0.05 wt. %, the additional catalyst was introduced at the level of 1/3 of the full height of the fast fluidized bed, and the additional catalyst was 5 wt. % catalyst circulation rate in the reactor. The distribution of the catalyst in the fast fluidized bed was set in a fully dense phase configuration by controlling the steam velocity and installing an umbrella-shaped catalyst distribution plate at the bottom of the fast fluidized bed such that the axial solids fraction ε of the catalyst in the fast fluidized bed varied within the range of 0 ,1 to 0.2 from bottom to top. The reaction product and the spent catalyst were rapidly separated, and the reaction product was further separated in a product separation system according to the distillation range. The spent catalyst was passed into the stripper by gravity, the hydrocarbon vapors that were adsorbed on the spent catalyst were separated by steam stripping, the stripper-cleaned catalyst was directly passed into the regenerator without heat exchange, and brought into contact with air for regeneration by coke burning, and the regenerated catalyst was recycled into the reactor for reuse. The process conditions used and the distribution of products are shown in Table 2.

As can be seen in Table 2, in this example, the ethylene yield is 5.2 wt. %, the yield of propylene is 18.2 wt. %, the yield of light aromatic compounds is 11.5 wt. %, and the yields of dry gas and coke are 10.8 wt. % and 8.5 wt. %, respectively.

Example 2

This example was carried out in accordance with the process flow diagram shown in figure 2 on a medium-sized plant, using a hydrogenated residue as a feed oil, a DMMC-2 catalyst as a catalyst, and a combined reactor containing a fluidized bed connected in series with transport in a rarefied phase and a fast fluidized bed as a reactor. The preheated feed oil was introduced into the bottom of the fluidized bed with transport in rarefied phase to contact the catalytic cracking catalyst for the first cracking reaction. The effluent from the reactor moved upward into the full dense phase fast fluidized bed reaction zone for mixing with additional regenerated catalyst for the second catalytic cracking reaction. The carbon content of the additional regenerated catalyst was 0.05 wt. %, the additional catalyst was introduced at the level of 1/3 of the full height of the fast fluidized bed, and the additional catalyst was 5 wt. % catalyst circulation rate in the reactor. The distribution of the catalyst in the fully dense phase reaction zone was set to the fully dense phase configuration by adjusting the steam rate and installing an umbrella-shaped catalyst distribution plate in the lower section of the fully dense phase reaction zone such that the axial solid fraction ε of the catalyst in the fully dense phase reaction zone varied within range from 0.1 to 0.2 from bottom to top. The reaction product and the spent catalyst were rapidly separated, and the reaction product was further separated in a product separation system according to a distillation range. The spent catalyst was passed into the stripper by gravity, the hydrocarbon vapors that were adsorbed on the spent catalyst were separated by steam stripping, the purified catalyst in the stripper was directly passed into the regenerator without heat exchange, and brought into contact with air for regeneration by coke burning, and the regenerated catalyst was recycled into the reactor for reuse. The mixed fraction C4 obtained after the separation of the reaction products was recycled to the lower part of the fluidized bed with transport in a rarefied phase for the subsequent reaction. The process conditions used and the distribution of products are shown in Table 2.

As can be seen in Table 2, in this example, the ethylene yield is 5.9 wt. %, the yield of propylene is 21.1 wt. %, the yield of light aromatic compounds is 11.8 wt. %, and the yields of dry gas and coke are 10.7 wt. % and 8.4 wt. %, respectively.

Comparative Example 1

This comparative example was carried out in a medium size plant using a hydrogenated residue as the feed oil, a DMMC-2 catalyst as a catalyst, and a combined reactor containing a vertical reactor and a fluidized bed connected in series as a reactor. The preheated feed oil was introduced into the bottom of the vertical reactor to contact the catalyst for the catalytic cracking reaction. The resulting reaction vaporous hydrocarbons, steam and spent catalyst were passed into the dense phase fluidized bed through the outlet of the vertical reactor for subsequent reaction. The reactor effluent was passed into a closed cyclone where the reaction product and the spent catalyst were quickly separated, and the reaction product was further separated in a product separation system according to the distillation range. The spent catalyst was passed into the stripper by gravity, the hydrocarbon vapors that were adsorbed on the spent catalyst were separated by steam stripping, the stripper-cleaned catalyst was directly passed into the regenerator without heat exchange, and brought into contact with air for regeneration by coke burning, and the regenerated catalyst was recycled into a vertical reactor for reuse. The process conditions used and the distribution of products are shown in Table 2.

As can be seen from the results of Table 2, in this comparative example, the yield of ethylene is 3.7 wt. %, the yield of propylene is 12.8 wt. %, the yield of light aromatic compounds is 5.5 wt. %, and the yields of dry gas and coke are 12.9 wt. % and 13.3 wt. %, respectively.

Comparative Example 2

This comparative example was essentially the same as example 1, except that no catalyst distribution plate was present at the bottom of the fast fluidized bed, and the axial solids fraction s of the catalyst in the fast fluidized bed exhibited a gradient of 0.1 → 0.2 → 0.3 from top to bottom. The process conditions used are the same as in Example 1 and the product distribution is shown in Table 2.

From the results of the above examples and comparative examples, it can be seen that the process according to the present application provides higher yields of ethylene, propylene and light aromatics, while providing reduced yields of dry gas and coke.

Preferred embodiments according to the present application have been described in detail above, but the present application is not limited to the details of the above embodiments. Various modifications of the technical solution according to this application can be made within the scope of the corresponding inventive concept, and these modifications should be within the scope of this application.

It should be noted that the various features described in the present embodiments may be combined in any suitable manner, as long as there is no conflict. To avoid unnecessary repetition, these possible combinations are not described in this application.

In addition, the various embodiments of the present application may also be combined arbitrarily, and these combinations should also be considered part of the present application, provided that they do not deviate from the spirit of the present application.

Claims (42)

1. The method of catalytic cracking of liquid petroleum products, including the stage of contacting the original liquid petroleum products in contact with a catalytic cracking catalyst for reaction in a reactor containing a fluidized bed with transport in a rarefied phase and with a fast fluidized bed connected in series, and in the fast fluidized bed axial solid fraction catalyst ε is measured at the top, middle and bottom of the reaction zone, divided into equal parts in the axial direction, in each case is not less than about 0.1 and not more than about 0.2, providing an axial solids fraction ε of the catalyst, which is controlled within the range from about 0.1 to about 0.2 from bottom to top, moreover, the axial solid fraction of the catalyst ε = pressure difference between two positions in the reaction zone along the axial direction, measured by a pressure difference sensor / distance between two positions along the axial direction / catalyst particle density, where the pressure drop is expressed in kg/m ² , the distance between two positions along the axial direction is expressed in meters (m), and the density of the catalyst particles is expressed in kg/m ³ ; catalyst particle density = theoretical density / (catalyst pore volume × theoretical density + 1), where the theoretical density is expressed in kg / m ³ and the pore volume of the catalyst is expressed in m ³ , and the theoretical density and pore volume of the catalyst are determined by the methods of pycnometric research and water titration, respectively, and the method is used to obtain light olefins from heavy starting oil products, and it additionally includes the following steps: i) contacting the first feedstock, comprising the light feedstock and/or the heavy feedstock, with a fluid catalytic cracking catalyst in a rarefied phase transport to effect a first catalytic cracking reaction; And ii) introducing the reactor effluent from step (i) and optionally a second feed containing light feed and/or heavy feed oil into a second FCC reaction under effective conditions to produce light olefins, wherein the axial solids fraction ε catalyst in a fast fluidized bed regulate within the range from about 0.1 to about 0.2 from bottom to top; iii) separating the effluent from the reactor to obtain a light olefin rich reaction product and a spent catalyst; iv) regenerating the spent catalyst and recycling at least a portion of the resulting regenerated catalyst to step (i) as a catalytic cracking catalyst; And v) optional separation of reaction products to produce dry gas, liquefied petroleum gas, gasoline, diesel fuel and catalyst slurry, wherein the light feedstock is selected from the group consisting of a C4 hydrocarbon fraction, a C5-C6 light gasoline fraction, and combinations thereof, and at least one of the first and second feedstocks contains a heavy feedstock, and said method also includes: introducing one or multiple additional catalyst streams into the fast fluidized bed and contacting the additional catalyst with the material in the fast fluidized bed to carry out the catalytic cracking reaction, moreover, a single or each of several streams of additional catalyst independently has a carbon content of approximately from 0 to 1.0 wt. %, and each is independently selected from regenerated, semi-regenerated, or spent catalytic cracking catalysts, and the total amount of one or more streams of additional catalyst is from about 5 to 50 wt. % relative to the catalyst circulation rate in the reactor; wherein the single or multiple additional catalyst streams are independently introduced at a level of about 20 to 90% of the total height of the fast fluidized bed. 2. The method according to p. 1, in which the total amount of one or more streams of additional catalyst is from about 5 to 30 wt. % with respect to the catalyst circulation rate in the reactor. 3. The process of claim 1, wherein the single or multiple additional catalyst streams are independently introduced at a level of from about 30% to about 75%. 4. The method according to any one of paragraphs. 1-3, in which the catalytic cracking catalyst contains, in relation to the dry weight of the catalytic cracking catalyst, from about 1 wt. % to about 50 wt. % zeolite, from about 5 wt. % to about 99 wt. % inorganic oxide, and from about 0 wt. % to about 70 wt. % clay; wherein the zeolite comprises a mesoporous zeolite selected from the group consisting of ZSM zeolites, ZRP zeolites and combinations thereof, and optionally a macroporous zeolite selected from the group consisting of rare earth-substituted Y-type zeolites, rare-earth-substituted H-Y type zeolites, ultrastable zeolites Y-type, high-silica Y-type zeolites, and combinations thereof. 5. The method according to any one of paragraphs. 1-4, in which the first catalytic cracking reaction is carried out under the following conditions: the reaction temperature is from about 500 to 600°C, the reaction time is from about 0.05 to 5 seconds, the mass ratio of catalyst and oil is from about 1:1 to about 50:1, water to oil weight ratio from about 0.03:1 to about 0.5:1, catalyst density from about 20 to 100 kg/m ³ , steam velocity from about 4 to 18 m/s , the reaction pressure is about 130 to 450 kPa, and the mass flow rate G _s of the catalyst flow is about 180 to 500 kg/(m ² s); And while the second catalytic cracking reaction is carried out under the following conditions: the reaction temperature is from about 510 to 650°C, the reaction time is from about 1 to 20 seconds, the mass ratio of catalyst and oil is from about 3:1 to about 50:1, mass water to oil ratio is about 0.03:1 to about 0.8:1, catalyst density is about 120 to 290 kg/m ³ , steam velocity is about 0.8 to 2.5 m/s, pressure reaction is about 130 to 450 kPa, and the mass flow rate G _s of the catalyst flow is about 15 to 150 kg/(m ² s). 6. The method according to any one of paragraphs. 1-5, wherein the heavy feedstock is a low grade heavy oil whose properties meet at least one of the following criteria: density at 20°C is about 900 to 1000 kg/m ³ , carbon residue is about 2 to 10 wt. %, the total content of nickel and vanadium is from about 2 to 30 ppm, and the characteristic factor K is less than about 12.1. 7. The method according to any one of paragraphs. 1-6, wherein the heavy oil feedstock is a low grade heavy oil product selected from heavy petroleum hydrocarbons, other mineral oils, and combinations thereof; wherein the heavy petroleum hydrocarbon is selected from the group consisting of vacuum distillation residue, low grade atmospheric distillation residue, low grade hydrogenated residue, coker gas oil, deasphalted oil, vacuum gas oil, high acid number crude oil, high metal crude oil, and combinations thereof. ; And wherein said other mineral oil is selected from the group consisting of coal liquefaction oil, oil sand bitumen, shale oil, and combinations thereof. 8. The method according to any one of paragraphs. 1-7, wherein at least one of the first and second feedstocks comprises a light feedstock selected from the group consisting of C4 hydrocarbon cut, C5-C6 light gasoline cut, and combinations thereof. 9. The method of claim 8, wherein the first feedstock comprises a light feedstock and a heavy feedstock, and at least a portion of the light feedstock is introduced into the fluidized bed with transport in rarefied phase at a position upstream of the position where the heavy feedstock is injected into a fluidized bed with transport in a rarefied phase; wherein the first feed contains a heavy feed oil and the second feed contains a light feed; or the first feedstock contains a light feedstock, and the second feedstock contains a heavy feedstock. 10. The method according to any one of paragraphs. 1-9, in which the fast fluidized bed consists of a bottom-up fully dense phase reaction zone and a transition section, and the fully dense phase reaction zone is present in the form of a hollow column having an approximately circular cross section with a constant diameter or a variable diameter, an open lower end and an open upper end, a fluidized bed with rarefied phase transport is in communication with the lower end of the fully dense phase reaction zone, the upper end of the fully dense phase reaction zone is in communication with the outlet section of the reactor through a transition section, the lower section of the fully dense phase reaction zone optionally contains one or more inlets for introduction of the second source material, wherein the cross-sectional diameter of the lower end of the fully dense phase reaction zone is greater than or the same as the diameter of the fluidized bed with transport in rarefied phase, and the cross-sectional diameter of the upper end of the fully dense phase reaction zone is greater than the diameter of the outlet section, and the bottom or side wall of the fully dense phase reaction zone contains one or more additional catalyst inlets, each independently located at 30 to 75% of the full height of the fast fluidized bed. 11. The process of claim 10 wherein the fully dense phase reaction zone is in the form of a hollow cylinder of constant diameter or a hollow column whose diameter increases continuously or discontinuously from bottom to top. 12. A system suitable for use in the catalytic cracking of liquid petroleum products, in particular heavy crude oils, comprising a catalytic cracking reactor, a catalyst separating device, an optional reaction product separator, and a regenerator, wherein the catalytic cracking reactor comprises a fluidized bed with a rarefied phase transport and a fast fluidized bed connected in series, wherein the fluidized bed with the rarefied phase transport is in fluid transfer communication with the fast fluidized bed and is located upstream of the relatively fast fluidized bed along direction of movement of the reaction stream; a fluidized bed with transport in a rarefied phase contains a catalyst inlet in the lower part and an inlet of the first source material in the lower part, a fast fluidized bed contains an outlet in the upper part and an optional inlet of the second source material in the lower part, the catalyst separating device contains an inlet, an outlet of the catalyst and an outlet reaction products, the optional reaction product separator comprises a reaction products inlet, a dry gas outlet, an LPG outlet, a gasoline outlet, a diesel fuel outlet, and a catalyst slurry outlet, and the regenerator comprises a catalyst inlet and a catalyst outlet, the fluidized bed catalyst inlet to the rarefied phase transport is in fluid transfer communication with the regenerator catalyst outlet, the fast fluidized bed outlet is in fluid transfer communication with the catalyst separating device inlet, the reaction products outlet of the catalyst separating device is in communication with the possibility of transferring fluid with the inlet of the reaction products of the optional separator of the reaction products, and the catalyst outlet of the catalyst separating device is in communication with the possibility of transferring fluid with the inlet of the catalyst of the regenerator, wherein the fast fluidized bed consists of a fully dense-phase reaction zone and a transition section arranged in series from bottom to top, and the fully dense-phase reaction zone is present in the form of a hollow column having an approximately circular cross section with a constant diameter or a variable diameter, an open lower end and an open upper end, a fluidized bed with transport in a rarefied phase is in communication with the lower end of the fully dense phase reaction zone, the upper end of the fully dense phase reaction zone is in communication with the outlet section of the reactor through the transition section, the lower section of the fully dense phase reaction zone optionally contains one or more inlets of the second source material, wherein the cross-sectional diameter of the lower end of the fully dense-phase reaction zone is greater than or the same as the diameter of the fluidized bed with transport in rarefied phase and the cross-sectional diameter of the upper end of the fully dense-phase reaction zone is greater than the diameter of the outlet section, and the lower or side the wall of the fully dense phase reaction zone contains one or more additional catalyst inlets, each independently located at about 20 to 90% of the full height of the fast fluidized bed. 13. The system of claim 12 wherein the single or multiple additional catalyst streams are independently located at a level of about 30 to 75% of the full height of the fast fluidized bed. 14. The system of claim 12 or 13 wherein the fast fluidized bed is positioned coaxially with the rarefied phase transport fluidized bed and the fast fluidized bed is positioned above the rarefied phase transport fluidized bed. 15. The system of claim 14 wherein preferably the rapid fluidized bed comprises a catalyst distributor plate at the bottom and/or a gas distributor is present at the second feed inlet position. 16. The system according to any one of paragraphs. 12-15, in which the fully dense phase reaction zone is present in the form of a hollow cylinder of constant diameter or a hollow column whose diameter increases continuously or discontinuously from bottom to top. 17. The system of claim 16, wherein the fully dense-phase reaction zone is in the form of an inverted truncated cone with a longitudinal section in the shape of an isosceles trapezoid, and has a lower cross-sectional diameter of approximately 0.2 to 10 meters; and the ratio of the diameter of the upper cross section and the diameter of the lower cross section, comprising from more than 1 to about 50; the ratio of the maximum cross-sectional diameter of the fully dense phase reaction zone to the total height of the fast reaction bed is from about 0.005:1 to about 1:1; the ratio of the height of the fully dense phase reaction zone to the total height of the fast reaction bed is from about 0.1:1 to about 0.9:1, and/or the transition section is in the form of a hollow truncated cone with a longitudinal section in the form of an isosceles trapezoid with sides having an internal inclination angle α of approximately 25 to 85°, and the ratio of the height of the transition section to the total height of the fast reaction layer is from approximately 0.1: 1 to about 0.9:1.

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