A method and system for catalytic cracking to produce light olefins and marine fuel oil

Abstract

The present invention relates to a method and system for producing low-carbon olefins and marine fuel oil by catalytic cracking, the method comprising: sending a heavy raw material into a downcomer reactor to contact with a catalyst for catalytic cracking reaction, sending a first mixed product into a settling section for gas-solid separation; sending a light raw material into a riser reactor to contact with a catalyst and carry out a catalytic cracking reaction, sending a second mixed product into a fluidized bed reactor to contact with a first semi-regenerated catalyst and a catalyst and carry out a catalytic cracking reaction, so that a third mixed product is separated into a third oil gas and a regenerated catalyst in the settling section; regenerating the regenerated catalyst to obtain a regenerated catalyst; separating the oil gas from the settling section in a separation device, sending the separated diesel oil into a hydro-reforming reactor for hydro-reforming, and sending the obtained hydro-reforming diesel oil into a riser reactor. The method of the present invention can produce more marine fuel oil and reduce the coke yield, and ensure the yield of low-carbon olefins.

Description

Technical Field

[0001] A method and system for producing light olefins and marine fuel oil by catalytic cracking. Background Art

[0002] Since the mid-1980s, the Petrochemical Research Institute of Sinopec has been engaged in the research of catalytic cracking family technology for producing light olefins from heavy oil as raw materials, and has successfully developed catalytic cracking (DCC, USP4980053 and USP5670037) technology for the maximum production of propylene and catalytic thermal cracking (CPP, USP6210562) for the maximum production of ethylene. So far, these two technologies mainly use a single riser reactor or a single riser reactor combined with a dense phase fluidized bed reactor structure, and the dry gas and coke yields are also high while improving the yield of light olefins. At present, the trend of heavy and inferior crude oil in my country is becoming more and more obvious, the conversion rate of catalytic cracking reaction of inferior heavy oil is low, the coke yield is high, and the quality of gasoline and diesel obtained is poor, which has seriously affected the benefits of the catalytic cracking process of inferior heavy oil.

[0003] For inferior raw materials, the cracking depth is reasonably controlled. On the one hand, the easily cracked hydrocarbons in the raw materials are cracked into high-molecular-weight olefins such as ethylene, propylene and butene as the most basic organic synthesis raw materials; on the other hand, the hydrocarbons that are not easy to crack are used as marine fuel oil. This can effectively avoid the shortcomings of high coke yield and poor quality of gasoline and diesel in the catalytic cracking reaction of inferior raw materials.

[0004] At present, the main production process of small molecule olefins in the world is steam cracking process, but high temperature cracking furnace is easy to coke, so this process basically uses light oil as raw material, such as natural gas, naphtha, light diesel, and hydrocracking tail oil as raw material. In recent years, the technology of integrating multiple processes to crack heavy oil to produce more low-carbon olefins, aromatics and fuel oil has received great attention. These technologies are to select different reactors, different processing processes and different catalysts for different raw materials, to ensure that various raw materials react in a reaction environment that is more suitable for their own characteristics to generate different products, and to ensure the maximization of production efficiency.

[0005] Chinese patent CN101074392A discloses a method for producing propylene and high-quality gasoline and diesel by two-stage catalytic cracking. The method mainly utilizes a two-stage riser catalytic process, adopts a catalyst rich in shape-selective molecular sieves, uses heavy petroleum hydrocarbons or various animal and plant oils rich in hydrocarbons as raw materials, optimizes the feeding method for reaction materials of different properties, and controls the appropriate reaction conditions for different materials to achieve the purpose of improving propylene yield, taking into account the light oil yield and quality, and inhibiting the formation of dry gas and coke. It specifically proposes that the first stage riser feed is fresh heavy crude oil, and light hydrocarbon feed can be fed into its lower part or bottom; the second stage riser feed is gasoline and circulating oil with high olefin content, which can be fed in layers or mixed, and other light hydrocarbon feed can be fed into its lower part or bottom.

[0006] Chinese patent CN101045667A proposes a catalytic conversion method for improving the yield of light olefins. In this method, hydrocarbon oil raw materials are injected into a downward reactor through a raw material nozzle, contacted with a regenerated catalyst and an optional carbon deposition catalyst, and cracked products and catalysts to be regenerated are separated. Light olefins are obtained after the cracked products are separated, and at least a portion of the remaining products are introduced into a riser reactor to contact and react with a regeneration agent to separate oil and gas from the catalyst to be regenerated. This method separates the generated light olefins from the regeneration agent in time, and strives to effectively suppress the secondary reaction of light olefins and improve the yield of light olefins. However, it is difficult to meet the conversion rate of heavy oil and light hydrocarbons with only a downward reactor and a riser reactor, and it is also impossible to maximize the yield of light olefins. Moreover, it can be seen from the embodiments of the patent that the ratio of light olefins to dry gas yield is below 3, the raw materials cannot be fully utilized, and the low-value products are high.

[0007] Chinese patent CN101210191A proposes a catalytic cracking method in which a descending reactor and a riser reactor are connected in series. The preheated crude oil enters the descending reactor and contacts with the high-temperature regenerated catalyst from the regenerator, vaporizes and undergoes a cracking reaction, the oil and gas from the outlet of the descending reactor enter the riser reactor to continue the reaction, another stream of regenerated catalyst is introduced from the inlet of the riser reactor, and the oil and gas from the outlet of the riser reactor and the catalyst enter the sedimentation separator for separation. Depending on the target product, a catalyst different from that of the descending reactor can be used in the riser reactor, which can increase the gasoline yield and improve the product quality. However, the light hydrocarbons are not further converted, so the yield of low-carbon olefins is not high.

[0008] Patent US6123830 discloses a combined process of two-stage catalytic cracking and two hydroprocessing. Its purpose is to produce olefins in the largest amount, improve the quality of distillate oil and the octane number of gasoline. The raw oil first passes through the first hydroprocessing unit to obtain the first hydroprocessing tail oil; the first hydroprocessing tail oil enters the first catalytic cracking unit, the catalyst active component of which is mainly a large-pore molecular sieve, to obtain naphtha, diesel and heavy oil, of which the heavy oil enters the second hydroprocessing unit for hydrogenation to obtain the second hydroprocessing tail oil, and the second hydroprocessing tail oil enters the second catalytic cracking unit for cracking, the catalyst active component of which is mainly a medium-pore molecular sieve, and the propylene yield of this method is low.

[0009] Chinese patent CN1721510A proposes a combined process for producing low-carbon olefins and aromatics. The raw oil is first hydrotreated, and the resulting hydrogenated tail oil is subjected to catalytic cracking reaction. The generated light hydrocarbons are recycled to the catalytic cracking unit, and LCO is returned to the hydrotreating unit. Naphtha and steam cracking naphtha are selectively hydrogenated and solvent extracted in turn to obtain raffinate oil and one of the target products, aromatics. The raffinate oil and hydrogenated naphtha are recycled to the steam cracking unit. This method maximizes the production of low-carbon olefins from heavy raw materials, wherein the yield of propylene exceeds 30% by weight, and aromatics such as toluene and xylene are co-produced. Summary of the invention

[0010] The object of the present invention is to provide a method and system for producing light olefins and marine fuel oil by catalytic cracking, which can produce more marine fuel oil and reduce the coke yield, thereby ensuring the yield of light olefins.

[0011] In order to achieve the above object, the first aspect of the present invention provides a method for producing light olefins and marine fuel oil by catalytic cracking, the method comprising:

[0012] S1, feeding the heavy raw material into the upper part of the descending tube reactor to contact with the first catalytic cracking catalyst from the top of the descending tube reactor and performing a first catalytic cracking reaction from top to bottom to obtain a first mixed product;

[0013] S2, sending the first mixed product to a settling section of a fluidized bed reactor for gas-solid separation to obtain a first oil gas and a first semi-spent catalyst, and drawing out the first oil gas from above the settling section;

[0014] S3, sending the light raw material into the lower part of the riser reactor to contact with the second catalytic cracking catalyst and perform a second catalytic cracking reaction from bottom to top to obtain a second mixed product;

[0015] S4, sending the second mixed product into a fluidized bed reactor to contact the first semi-spent catalyst and the third catalytic cracking catalyst and perform a third catalytic cracking reaction to obtain a third mixed product; separating the third mixed product into a third oil gas and a spent catalyst in the settling section; drawing out the third oil gas from the top of the settling section; sending the spent catalyst into a catalyst regenerator for regeneration to obtain a regenerated catalyst;

[0016] S5. Send the first oil and gas and the third oil and gas from the settling section to a separation device for separation, and send the separated diesel to a hydro-modification reactor to contact with a hydro-catalyst for hydro-modification; send the obtained hydro-modified diesel to the riser reactor for the second catalytic cracking reaction.

[0017] Optionally, the conditions of the first catalytic cracking reaction include: a temperature of 420-600°C, a catalyst-oil ratio of 2-14, and a reaction time of 0.2-4 seconds;

[0018] The conditions of the second catalytic cracking reaction include: temperature of 480-699°C, catalyst-oil ratio of 5-23, and reaction time of 1-10 seconds;

[0019] The conditions of the third catalytic cracking reaction include: a temperature of 460-640°C, a weight hourly space velocity of 1-35 hours -1 , the reaction pressure is 0.15-0.35 MPa;

[0020] The conditions of the hydrogenation reforming reaction include: hydrogen partial pressure of 1.2-8.0 MPa, reaction temperature of 150-300°C, hydrogen-oil volume ratio of 150-300, volume space velocity of 1-20h -1 .

[0021] Optionally, in step S2, the first mixed product is fed into the middle part of the settling section of the fluidized bed reactor for the gas-solid separation.

[0022] Optionally, the weight ratio of the light raw material fed into the riser reactor to the heavy raw material fed into the downcomer reactor is (0.05-0.30):1, preferably (0.08-0.20):1;

[0023] The weight ratio of the hydrotreated and reformed diesel fed into the riser reactor to the heavy feedstock fed into the downcomer reactor is (0.02-0.30):1, preferably (0.05-0.20):1.

[0024] Optionally, step S4 further comprises: feeding part of the regenerated catalyst from the catalyst regenerator as the first catalytic cracking catalyst into the top of the downcomer reactor; feeding the remaining part of the regenerated catalyst as the second catalytic cracking catalyst and the third catalytic cracking catalyst into the fluidized reaction section of the riser reactor and the fluidized bed reactor respectively;

[0025] Preferably, based on the total weight of the regenerated catalyst leaving the catalyst regenerator per unit time, 10-70 wt% of the regenerated catalyst is fed into the downcomer reactor, 0-30 wt% of the regenerated catalyst is fed into the fluidized reaction section of the fluidized bed reactor, and 10-40 wt% of the regenerated catalyst is fed into the riser reactor.

[0026] Optionally, step S3 further comprises: first sending the catalyst to be regenerated into the stripping section of the fluidized bed reactor for stripping, and then sending the stripped catalyst to be regenerated into the catalyst regenerator for regeneration.

[0027] Optionally, the regenerated catalyst contains molecular sieve, inorganic oxide and optional clay; based on the dry weight of the regenerated catalyst, the content of the molecular sieve is 2-51% by weight, the content of the inorganic oxide is 6-97% by weight, and the content of the clay is 5-75% by weight;

[0028] The molecular sieves include a shape-selective molecular sieve with an average pore size of less than 0.7 nanometers and a Y-type molecular sieve; on a dry basis and based on the total weight of the molecular sieves, the content of the shape-selective molecular sieve with an average pore size of less than 0.7 nanometers is 45-95% by weight, and the content of the Y-type molecular sieve is 2-54% by weight;

[0029] The shape-selective molecular sieve with an average pore size of less than 0.7 nanometers is selected from one or more of the ZSM series molecular sieves, ZRP molecular sieves, ferrierite, chabazite, cyclopentane, erionite, A zeolite, column zeolite and lanoite, and the Y-type molecular sieve is selected from one or more of the rare earth Y-type molecular sieve, rare earth hydrogen Y-type molecular sieve, ultra-stable Y-type molecular sieve and rare earth ultra-stable Y-type molecular sieve.

[0030] Optionally, the hydrogenation catalyst contains a carrier and an active metal component; the content of the active metal component is 1-20 wt% based on the dry weight of the hydrogenation catalyst;

[0031] The carrier is selected from alumina and / or amorphous silica-alumina, and the active metal component is selected from non-noble metals of Group VIB and / or non-noble metals of Group VIII.

[0032] Optionally, in step S1, the heavy raw material is one or more selected from vacuum wax oil, atmospheric residue oil, coker wax oil, deasphalted oil, furfural refined raffinate oil, coal liquefaction oil, oil sand oil and shale oil;

[0033] In step S3, the light feedstock is an olefin-rich liquefied gas component and / or an olefin-rich light gasoline component; the hydrotreated and reformed diesel fed into the riser reactor for the second catalytic cracking reaction is part or all of the heavy fraction from the hydrotreatment and reforming reactor.

[0034] A second aspect of the present invention provides a system for producing light olefins and marine fuel oil, the system comprising a downcomer reactor, a fluidized bed reactor, a catalyst regenerator, a riser reactor, a separation device and a hydro-reforming reactor;

[0035] The down-tube reactor is provided with a catalyst inlet at the top, a heavy raw material inlet at the upper part and a product outlet at the bottom; the fluidized bed reactor comprises a settling section, a first material inlet at the settling section, a gas outlet above the settling section, a second material inlet below the settling section, a catalyst inlet and a catalyst outlet; the catalyst regenerator is provided with a catalyst inlet and a catalyst outlet; the riser reactor is provided with a catalyst inlet at the bottom, a raw material inlet at the lower part and a product outlet at the top; the separation device is provided with a material inlet and a diesel outlet, and the hydro-reforming reactor is provided with a material inlet and a product outlet;

[0036] The product outlet of the downtube reactor is connected to the first material inlet of the fluidized bed reactor, the product outlet of the riser reactor is connected to the second material inlet of the fluidized bed reactor, the catalyst inlet of the catalyst regenerator is connected to the catalyst outlet of the fluidized bed reactor, the gas outlet of the fluidized bed reactor is connected to the inlet of the separation device, the diesel outlet of the separation device is connected to the material inlet of the hydro-modification reactor, and the product outlet of the hydro-modification reactor is connected to the raw material inlet of the riser reactor.

[0037] Optionally, the catalyst inlet of the fluidized bed reactor, the catalyst inlet of the downcomer reactor and the catalyst inlet of the riser reactor are each independently connected to the catalyst outlet of the catalyst regenerator.

[0038] Optionally, the fluidized bed reactor further comprises a stripping section fluidly connected to the fluidized reaction section, the stripping section is arranged downstream of the fluidized reaction section, and a catalyst outlet of the fluidized bed reactor is arranged at the lower part of the stripping section.

[0039] The method of the present invention adopts a reaction system combining a downtube reactor, a fluidized bed reactor, a riser reactor and a hydrogenation reforming reactor. By treating specific reaction raw materials in different reactors, more marine fuel oil can be produced and the coke yield can be reduced, the yield of light olefins can be guaranteed, and the ratio of the light olefin yield to the coke yield can be increased.

[0040] Other features and advantages of the present invention will be described in detail in the following detailed description. BRIEF DESCRIPTION OF THE DRAWINGS

[0041] The accompanying drawings are used to provide a further understanding of the present invention and constitute a part of the specification. Together with the following specific embodiments, they are used to explain the present invention but do not constitute a limitation of the present invention. In the accompanying drawings:

[0042] Figure 1 It is a structural schematic diagram of a specific implementation method of the present invention for catalytic cracking to produce light olefins and marine fuel oil.

[0043] Description of Reference Numerals

[0044] 1 Catalyst tank 2 Downstream tube reactor 3 Fluidized bed reactor

[0045] 4 Sedimentation section 5 Stripping section 6 Waiting for production inclined tube

[0046] 7 Catalyst regenerator 8 Regeneration inclined tube 9 Riser reactor

[0047] 10 Regeneration inclined tube 11 Regeneration inclined tube 12 Separation device

[0048] 13 Hydrogenation reactor 14 Hydrogenation product separation device 15 Pipeline

[0049] 16 pipeline 17 pipeline 18 pipeline

[0050] Pipeline 19 Pipeline 20 Pipeline 21

[0051] 22 pipeline 23 pipeline 24 pipeline

[0052] 25 pipeline 26 pipeline 27 pipeline

[0053] 28 pipeline 29 pipeline 30 pipeline

[0054] 31 pipeline 32 pipeline 33 pipeline

[0055] 34 pipeline DETAILED DESCRIPTION

[0056] The specific implementation of the present invention is described in detail below in conjunction with the accompanying drawings. It should be understood that the specific implementation described here is only used to illustrate and explain the present invention, and is not used to limit the present invention.

[0057] In the present invention, unless otherwise specified, directional words such as "up, down, left, right" generally refer to "up, down, left, right" when the system of the present invention can operate normally.

[0058] The first aspect of the present invention provides a method for producing light olefins and marine fuel oil by catalytic cracking, the method comprising:

[0059] S1, feeding the heavy raw material into the upper part of the descending tube reactor 2 to contact with the first catalytic cracking catalyst from the top of the descending tube reactor 2 and performing a first catalytic cracking reaction from top to bottom to obtain a first mixed product;

[0060] S2, sending the first mixed product to the settling section 4 of the fluidized bed reactor 3 for gas-solid separation to obtain first oil and gas and a first semi-spent catalyst, and drawing out the first oil and gas from the top of the settling section 4;

[0061] S3, feeding the light raw material into the lower part of the riser reactor 9 to contact with the second catalytic cracking catalyst and perform a second catalytic cracking reaction from bottom to top to obtain a second mixed product;

[0062] S4, sending the second mixed product into a fluidized bed reactor 3 to contact the first semi-spent catalyst and the third catalytic cracking catalyst and perform a third catalytic cracking reaction to obtain a third mixed product; separating the third mixed product into a third oil gas and a spent catalyst in the settling section 4; drawing out the third oil gas from the top of the settling section 4; sending the spent catalyst into a catalyst regenerator 7 for regeneration to obtain a regenerated catalyst;

[0063] S5. Send the first oil and gas and the third oil and gas from the settling section 4 to the separation device 12 for separation, and send the separated diesel to the hydrogenation reactor 13 to contact with the hydrogenation catalyst for hydrogenation; send the obtained hydrogenated diesel to the riser reactor 9 for the second catalytic cracking reaction.

[0064] The method of the present invention adopts a reaction system combining a downtube reactor, a fluidized bed reactor, a riser reactor and a hydro-reforming reactor, catalytically cracks the light raw materials and hydro-reforming diesel through the riser reactor and the fluidized bed reactor, and catalytically cracks the heavy raw materials through the downtube reactor, so that different raw materials can be catalytically cracked separately to improve the selectivity of each target product. And after the first catalytic cracking reaction product from the downtube reactor is introduced into the settling section of the fluidized bed reactor for gas-solid separation, the first semi-spent catalyst still having reaction activity obtained by separation is contacted and reacted with the second mixed product from the riser reactor, thereby improving the average activity of the catalyst. The marine fuel oil component directly leaves the reactor and no longer reacts, which can ensure that the marine fuel oil is retained to the maximum extent, optimize the reaction path, increase the production of light olefins while improving the yield of marine fuel components, reduce the yield of coke, gasoline and diesel, and improve the ratio of carbon olefin yield to coke yield.

[0065] In the method of the present invention, the generated diesel is further subjected to hydrogenation and reforming in a hydrogenation and reforming reactor to increase the hydrogen content of the diesel and saturate the polycyclic aromatic hydrocarbons. The riser reactor and the hydrogenation and reforming reactor are coupled and used, and the hydrogenated and reformed diesel is introduced into the riser reactor together with light hydrocarbons as a raw material. The two are in contact with a highly active catalyst in the riser reactor, and catalytic cracking reactions are more likely to occur, which can effectively improve the conversion rate of the raw material and the propylene yield.

[0066] According to the present invention, the catalytic cracking reaction is well known to those skilled in the art and will not be described in detail herein. In a specific embodiment of the present invention, the conditions of the first catalytic cracking reaction include: a temperature of 420-600°C, preferably 440-580°C, a catalyst-oil ratio of 2-14, preferably 7-12, a reaction time of 0.2-4 seconds, preferably 0.6-3 seconds, and a heavy raw material feed atomized water vapor accounting for 2-50% by weight of the sum of the weight of the heavy raw material and the atomized water vapor, preferably 5-15% by weight, wherein the temperature refers to the temperature at the bottom outlet of the down tube reactor.

[0067] In a specific embodiment of the present invention, the conditions of the second catalytic cracking reaction include: a temperature of 480-699°C, preferably 490-650°C, a catalyst-oil ratio of 5-23, preferably 6-20, a reaction time of 1-10 seconds, preferably 2-7 seconds, and the atomized water vapor of the light raw material feed accounts for 2-50% by weight of the sum of the weight of the light raw material, oil slurry and atomized water vapor, preferably 5-15% by weight, wherein the temperature refers to the temperature at the top outlet of the riser reactor.

[0068] In a specific embodiment of the present invention, the conditions of the third catalytic cracking reaction include: a temperature of 460-640°C, preferably 500-600°C, a weight hourly space velocity of 1-35 hours-1 , preferably 2-33 hours -1 , the reaction pressure is 0.15-0.35 MPa, preferably 0.2-0.35 MPa, wherein the reaction pressure refers to the absolute pressure at the outlet of the sedimentation section.

[0069] According to the present invention, the hydro-reforming reaction is also well known to those skilled in the art and will not be described in detail herein. In a specific embodiment of the present invention, the conditions of the hydro-reforming reaction include: a hydrogen partial pressure of 1.2-8.0 MPa, a reaction temperature of 150-300°C, a hydrogen-to-oil volume ratio of 150-300, and a volume space velocity of 1-20 h -1 Preferably, the hydrogen partial pressure is 1.6-6 MPa, the reaction temperature is 160-280°C, the hydrogen-to-oil volume ratio is 180-290 seconds, and the volume space velocity is 2-18h -1 The saturation of the hydromodified diesel condensed aromatics obtained under the above hydromodification conditions is higher. Introducing it into the riser reactor to carry out catalytic cracking reaction together with the light raw material is beneficial to further increase the yield of light olefins and reduce the yield of coke.

[0070] In a preferred embodiment of the present invention, in step S2, the first mixed product is sent to the middle part of the settling section 4 of the fluidized bed reactor 3 for the gas-solid separation to further produce more light olefins and marine fuel oil.

[0071] According to the present invention, in step S5, the product separated by the separation device 12 may also include: one or more of gaseous hydrocarbons, light gasoline, heavy gasoline and oil slurry. The separation device is commonly used by technicians in this field, for example, it can be a distillation tower.

[0072] According to the present invention, the hydrotreated and modified diesel fed into the riser reactor may be part or all of the hydrotreated and modified diesel. In a preferred embodiment of the present invention, the weight ratio of the hydrotreated and modified diesel fed into the riser reactor to the heavy feedstock fed into the descender reactor is (0.02-0.30):1, preferably (0.05-0.20):1.

[0073] In a preferred embodiment of the present invention, the weight ratio of the light feedstock fed into the riser reactor to the heavy feedstock fed into the downcomer reactor is (0.05-0.30):1, preferably (0.08-0.20):1.

[0074] According to the present invention, in a specific embodiment of the present invention, step S4 further comprises: sending part of the regenerated catalyst from the catalyst regenerator 7 as the first catalytic cracking catalyst to the top of the downcomer reactor 2; sending the remaining part of the regenerated catalyst as the second catalytic cracking catalyst and the third catalytic cracking catalyst to the riser reactor 9 and the fluidized bed reactor 3 respectively. Preferably, in order to promote the catalytic cracking reaction, effectively control the conversion rate, and improve the yield of light olefins in the riser reactor and the fluidized bed reactor, the regenerated catalysts sent to each reactor as the first catalytic cracking catalyst, the second catalytic cracking reaction and the third catalytic cracking reaction are all uncooled catalysts. In one embodiment, the temperature of the regenerated catalyst is 500-900°C, preferably 600-800°C.

[0075] In a preferred specific embodiment of the present invention, based on the total weight of the regenerated catalyst leaving the catalyst regenerator 7 per unit time, 10-70 wt% of the regenerated catalyst is fed into the downcomer reactor 2, 0-30 wt% of the regenerated catalyst is fed into the fluidized bed reactor 3, and 10-40 wt% of the regenerated catalyst is fed into the riser reactor 9. Preferably, 20-65 wt% of the regenerated catalyst is fed into the downcomer reactor 2, 8-28 wt% of the regenerated catalyst is fed into the fluidized bed reactor 3, and 15-30 wt% of the regenerated catalyst is fed into the riser reactor 9.

[0076] In a specific embodiment of the present invention, step S3 further includes: first sending the catalyst to be regenerated into the stripping section 5 of the fluidized bed reactor for stripping, and then sending the stripped catalyst to be regenerated into the catalyst regenerator 7 for the regeneration. With the above method, the water vapor stripped in the stripping section and the obtained oil and gas are simultaneously introduced into the bottom of the fluidized reaction section of the fluidized bed reactor, and are discharged from the reactor after passing through the fluidized reaction section, which can reduce the partial pressure of oil and gas, shorten the residence time of oil and gas in the settling section, increase the selectivity of propylene, and at the same time, after stripping, it is more conducive to restore the activity of the catalyst in the subsequent regeneration process.

[0077] According to the present invention, the catalyst used for catalytic cracking reaction is well known to those skilled in the art. In a specific embodiment of the present invention, the regenerated catalyst contains molecular sieve, inorganic oxide and optional clay; based on the dry weight of the regenerated catalyst, the content of the molecular sieve is 1-50% by weight, the content of the inorganic oxide is 5-99% by weight, and the content of the clay is 0-70% by weight. Preferably, in order to improve the yield of light olefins and increase the conversion rate of light raw materials and diesel, the molecular sieve can include a Y-type molecular sieve and a shape-selective molecular sieve with an average pore size of less than 0.7 nanometers; based on the dry basis and based on the total weight of the molecular sieve, the content of the shape-selective molecular sieve with an average pore size of less than 0.7 nanometers is 25-90% by weight, preferably 40-60% by weight, and the content of the Y-type molecular sieve is 10-75% by weight, preferably 30-65% by weight.

[0078] According to the present invention, shape-selective molecular sieves and Y-type molecular sieves are well known to those skilled in the art, and the shape-selective molecular sieves with an average pore size less than 0.7 nanometers may include but are not limited to one or more of ZSM series molecular sieves, ZRP molecular sieves, ferrierite, chabazite, cyclocrystal, erionite, A zeolite, column zeolite and turbidite, wherein the ZSM series molecular sieves may be selected from one or two of ZSM-5, ZSM-8, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35, ZSM-38, ZSM-48 and other molecular sieves of similar structure. The Y-type molecular sieve may include but are not limited to one or more of rare earth Y-type molecular sieves (REY), rare earth hydrogen Y-type molecular sieves (REHY), ultrastable Y-type molecular sieves (USY) and rare earth ultrastable Y-type molecular sieves (REUSY). The inorganic oxide, as a binder, may be silicon dioxide (SiO2) and / or aluminum oxide (Al2O3). The clay selected as the matrix, ie, the carrier, may be kaolin and / or halloysite.

[0079] In a specific embodiment of the present invention, the hydrogenation catalyst contains a carrier and an active metal component; based on the dry weight of the hydrogenation catalyst, the content of the active metal component is 1-20% by weight, preferably 3-15% by weight; the carrier is selected from alumina and / or amorphous silica-alumina, and the active metal component is selected from non-precious metals of Group VIB and / or non-precious metals of Group VIII.

[0080] According to the present invention, the heavy raw material is well known to those skilled in the art, for example, the heavy raw material is a heavy hydrocarbon raw material, and the heavy hydrocarbon can be one or more selected from petroleum hydrocarbons, mineral oils and synthetic oils. Petroleum hydrocarbons can be vacuum wax oil, atmospheric residue oil, vacuum wax oil mixed with part of vacuum residue oil or other hydrocarbon oils obtained by secondary processing, and the hydrocarbon oil obtained by secondary processing is one or more of coker wax oil, deasphalted oil, furfural refined raffinate oil. Mineral oil can be one or more selected from coal liquefaction oil, oil sand oil and shale oil. Synthetic oil can be a distillate oil obtained by FT synthesis of coal, natural gas or asphalt. In a specific embodiment of the present invention, in step S1, the heavy raw material is one or more selected from vacuum wax oil, atmospheric residue oil, coker wax oil, deasphalted oil, furfural refined raffinate oil, coal liquefaction oil, oil sand oil and shale oil.

[0081] According to the present invention, the light raw material introduced into the riser reactor is preferably olefin-rich gasoline and / or C4 hydrocarbons, and the olefin-rich gasoline is selected from the gasoline fraction produced by this process and / or the gasoline fraction produced by other devices. The gasoline fraction produced by other devices can be selected from one or more of catalytic cracking crude gasoline, catalytic cracking stabilized gasoline, coking gasoline, visbreaking gasoline and gasoline fractions produced by other oil refining or chemical processes, and the light gasoline fraction produced by this process is preferably selected. The olefin content of the olefin-rich gasoline can be 25-95% by weight, preferably 35-90% by weight, and preferably more than 50% by weight. The C4 hydrocarbon refers to a low molecular hydrocarbon compound that exists in the form of gas at room temperature and pressure with C4 fraction as the main component, including alkanes, alkenes and alkynes with a carbon number of 4. It can be a gaseous hydrocarbon product rich in C4 fraction produced by this process, or a gaseous hydrocarbon rich in C4 fraction produced by other device processes, wherein the C4 fraction produced by this process is preferred. The olefin content of the C4 hydrocarbons is greater than 50% by weight, preferably greater than 60% by weight, and most preferably greater than 70% by weight. In the light feedstock, the weight ratio of C4 hydrocarbons to gasoline can be (0-2):1, preferably (0-1.2):1, and more preferably (0-0.8):1. The weight ratio of the olefin-rich gasoline introduced into the riser reactor to the heavy feedstock introduced into the downcomer reactor can be (0.05-0.30):1, preferably (0.08-0.20):1.

[0082] In a specific embodiment of the present invention, in step S3, the light feedstock is an olefin-rich liquefied gas component and / or an olefin-rich light gasoline component; the hydrotreated and reformed diesel is part or all of the heavy fraction from the hydrotreatment and reforming reactor.

[0083] like Figure 1As shown, the second aspect of the present invention provides a system for producing light olefins and marine fuel oil, the system comprising a downcomer reactor 2, a fluidized bed reactor 3, a catalyst regenerator 7, a riser reactor 9, a separation device 12 and a hydro-reforming reactor 13;

[0084] The downcomer reactor 2 is provided with a catalyst inlet at the top, a heavy raw material inlet at the upper part and a product outlet at the bottom; the fluidized bed reactor 3 includes a settling section 4, a first material inlet at the settling section, a gas outlet above the settling section, a second material inlet below the settling section, a catalyst inlet and a catalyst outlet; the catalyst regenerator 7 is provided with a catalyst inlet and a catalyst outlet; the riser reactor 9 is provided with a catalyst inlet at the bottom, a raw material inlet at the lower part and a product outlet at the top; the separation device 12 is provided with a material inlet and a diesel outlet, and the hydro-reforming reactor 13 is provided with a material inlet, a liquid product outlet and a gas product outlet;

[0085] The product outlet of the downtube reactor 2 is connected to the first material inlet of the fluidized bed reactor 3, the product outlet of the riser reactor 9 is connected to the second material inlet of the fluidized bed reactor 3, the catalyst inlet of the catalyst regenerator 7 is connected to the catalyst outlet of the fluidized bed reactor 3, the gas outlet of the fluidized bed reactor 3 is connected to the inlet of the separation device 12, the diesel outlet of the separation device 12 is connected to the material inlet of the hydro-modification reactor 13, and the liquid product outlet of the hydro-modification reactor 13 is connected to the raw material inlet of the riser reactor 9.

[0086] In a specific embodiment of the present invention, the catalyst inlet of the fluidized bed reactor 3 , the catalyst inlet of the downcomer reactor 2 , and the catalyst inlet of the riser reactor 9 are each independently connected to the catalyst outlet of the catalyst regenerator 7 .

[0087] In a specific embodiment of the present invention, the fluidized bed reactor 3 also includes a stripping section 5 connected to the fluid of the fluidized reaction section, the stripping section 5 is arranged downstream of the fluidized reaction section, and the catalyst outlet of the fluidized bed reactor 3 is arranged at the lower part of the stripping section 5.

[0088] In a specific embodiment of the present invention, the system also includes a hydrogenation product separation device 14, which includes a raw material inlet, a gas outlet and a liquid outlet. The raw material inlet of the hydrogenation product separation device is connected to the product outlet of the hydrogenation reforming reactor 13, and the liquid outlet of the hydrogenation product separation device 14 is connected to the raw material inlet of the riser reactor 9.

[0089] The method provided by the present invention is further described below in conjunction with the accompanying drawings:

[0090] like Figure 1 As shown, the preheated heavy feedstock is mixed with the atomized water vapor from pipeline 15 in a certain proportion through pipeline 16, and then injected into the downtube reactor 2 to contact with the high-temperature first catalytic cracking catalyst (optionally part of it comes from the catalyst tank 1) and perform the first catalytic cracking reaction from top to bottom. The first mixed product obtained enters the settling section 4 through the outlet of the downtube reactor 2 for gas-solid separation of oil and gas and the catalyst. The first semi-generated catalyst and the first oil and gas obtained by the gas-solid separation, wherein the first semi-generated catalyst enters the fluidized reaction section of the fluidized bed reactor 3, and the first oil and gas enter the subsequent separation device 12 through pipeline 24. In the separation device 12, gaseous hydrocarbons (drawn out from pipeline 25), light gasoline (drawn out from pipeline 26), heavy gasoline (drawn out from pipeline 27), diesel (drawn out from pipeline 28) and slurry oil (drawn out from pipeline 29) are separated. The heavy gasoline drawn out from pipeline 27 can obtain propylene and a C4 fraction rich in olefins in subsequent product separation. The C4 fraction rich in olefins can be returned to the riser reactor 9 as a light feedstock and then converted into ethylene and propylene. The light gasoline drawn out from pipeline 26 can be used as a light feedstock to be converted into ethylene and propylene. The oil can be partially or completely returned to the reaction system for re-conversion; the diesel oil drawn out from pipeline 28 can be partially or completely entered into the hydro-modification reactor 13 for hydro-modification, and the hydrogen required for hydro-modification is introduced through pipeline 31; the product obtained by hydro-modification is sent to the hydrogenation product separator 14 through pipeline 32 for product separation, and is separated into hydrogen (drawn out from pipeline 33) and hydro-modified diesel oil (drawn out from pipeline 34), and the hydro-modified diesel oil is introduced into the riser reactor 9 for the second catalytic cracking reaction.

[0091] The preheated or non-preheated light hydrocarbon feedstock (for example, the olefin-rich gasoline fraction and / or C4 hydrocarbons from pipeline 26) and the hydrogenated modified diesel from the hydrogenation product separation unit 14 are mixed with the atomized water vapor from pipeline 20 in a certain proportion via pipeline 21, and then injected into the riser reactor 9, mixed with the high-temperature catalyst lifted by the lifting gas from pipeline 20 through the regeneration inclined tube 8, and subjected to a second catalytic cracking reaction to obtain a second mixed product. The second mixed product is introduced into the fluidized reaction section of the fluidized bed reactor 3 through the outlet of the riser reactor 9, contacts with the first semi-spent catalyst and the third catalytic cracking catalyst, and undergoes a third catalytic cracking reaction to obtain a third mixed product. The third mixed product is separated into third oil and gas and the catalyst to be regenerated in the settling section 4; the third oil and gas are led out of the fluidized bed reactor through pipeline 24 and enter the subsequent separation device 12, the catalyst to be regenerated enters the stripping section 5, the stripping steam is injected through pipeline 19, countercurrently contacts with the catalyst to be regenerated, and the reaction oil and gas carried by the catalyst to be regenerated is stripped out as much as possible, and then enters the catalyst regenerator 7 through the inclined pipe 6 to be regenerated for charring regeneration, the oxygen-containing gas is injected into the catalyst regenerator 7 through pipeline 18, the regenerated flue gas is led out through pipeline 23, and the regenerated catalyst is respectively introduced into the riser reactor 9, the downcomer reactor 2 and the fluidized bed reactor 3 through the regeneration inclined pipe 8, the regeneration inclined pipe 10 and the regeneration inclined pipe 11 for recycling.

[0092] In the above specific implementation process, the pre-lifting medium introduced into the pipeline 20 of the pre-lifting section of the riser can be selected from water vapor, C1-C4 hydrocarbons or conventional catalytic cracking dry gas, and water vapor is preferred in the present invention. The lifting medium introduced into the pipeline 17 to lift the regenerated catalyst into the catalyst tank 1 can be selected from water vapor, C1-C4 hydrocarbons, N2 or conventional catalytic cracking dry gas, and dry gas is preferred in the present invention.

[0093] The present invention is further illustrated by the following examples, but the present invention is not limited thereto.

[0094] The raw material oil and catalyst used in the embodiment and the comparative example are the same. The raw material A used is a cracking raw material, and its specific properties are shown in Table 1. The catalyst used is MMC-2 produced by Sinopec Qilu Catalyst Plant, containing a type-selective zeolite with an average pore size of less than 0.7 nanometers and a Y-type molecular sieve, and its specific properties are shown in Table 2.

[0095] Example 1

[0096] The test is Figure 1 The process is carried out in the system shown, which includes: a downcomer reactor, a fluidized bed reactor, a catalyst regenerator, a riser reactor, a separation device and a hydrogenation reforming reactor, wherein the fluidized bed reactor includes, from top to bottom, a fluid-connected settling section, a fluidized reaction section and a stripping section.

[0097] The product outlet at the bottom of the downtube reactor is connected to the middle fluid of the settling section. The catalytic cracking catalyst in the downtube reactor is a type B catalyst. The heavy raw material shown in Table 1 is subjected to a first catalytic cracking reaction. The first mixed product obtained by the reaction is separated in the settling section, and the first oil and gas obtained enters a separation device for separation, and the first semi-spent catalyst enters the fluidized reaction section of the fluidized bed reactor.

[0098] A fluidized bed reactor is connected in series at the top of the riser, and the catalytic cracking catalyst used in the descending tube reactor is a B-type catalyst. The light gasoline rich in olefins from the product separation device (distillation range is 30-85°C, olefin content is 52%, and weight accounts for 10% of the heavy raw materials introduced into the descending tube reactor) and the hydro-modified diesel from the hydro-modified reactor (distillation range is 204-350°C, and weight accounts for 7% of the heavy raw materials introduced into the descending tube reactor) are subjected to a second catalytic cracking reaction. The obtained second mixed product enters the fluidized reaction section of the fluidized bed reactor and contacts with the first semi-spent catalyst and the regenerated B-type catalyst. The third catalytic cracking reaction is carried out, and the third mixed product obtained enters the settling section for separation. The separated catalyst to be regenerated enters the stripping section for stripping and then enters the catalyst regenerator for regeneration. The obtained regenerated agent is again sent to the fluidized reaction section of the riser reactor, the descender reactor and the fluidized bed reactor for reaction. The third oil and gas obtained enters the separation device and is separated into gaseous hydrocarbons, light gasoline, heavy gasoline, diesel and oil slurry. Part of the diesel enters the hydro-reforming device for hydro-reforming reaction. The product enters the hydrogenation product separator and is separated into hydrogen and hydro-reformed diesel. The hydro-reformed diesel enters the riser for reaction again. The reaction operating conditions and reaction results are shown in Tables 3 and 4.

[0099] Example 2

[0100] Referring to Example 1, the only difference is that the weight of the hydro-modified diesel accounts for 12% of the heavy raw material introduced into the descending tube reactor, and the weight of the olefin-rich light gasoline accounts for 14% of the heavy raw material introduced into the descending tube reactor. The reaction operating conditions and reaction results are shown in Tables 3 and 4.

[0101] Comparative Example 1

[0102] The test was conducted in a medium-sized catalytic cracking unit, which includes an independent reaction and regeneration system: the reactor is a riser reactor plus a fluidized bed combined reactor, the top outlet of the riser reactor is connected in series with a fluidized bed reactor, and the raw materials shown in Table 1 are cracked; the oil and gas after the reaction are separated from the catalyst, the catalyst enters the stripper for stripping, enters the regenerator for regeneration, and then enters the riser reactor again for reaction; the reaction product from the fluidized bed reactor enters the product separation device. The reaction operating conditions and reaction results are shown in Tables 3 and 4.

[0103] Comparative Example 2

[0104] The experiment was conducted in the same manner as in Example 1, except that the hydrogenated reformed diesel was no longer introduced into the riser reactor, and the feedstock in the riser reactor was only olefin-rich light gasoline.

[0105] Comparative Example 3

[0106] The experiment was carried out in the same manner as in Example 1, with the only difference being that the first mixed product obtained from the downcomer reactor was not subjected to gas-solid separation but was directly introduced into the fluidized reaction section of the fluidized bed reactor, so that the first mixed product, the second mixed product and the regenerated type B catalyst in the fluidized reaction section were contacted to carry out catalytic cracking reaction.

[0107] Comparative Example 4

[0108] The experiment was carried out in the same manner as in Example 1, with the only difference being that the first mixed product obtained from the downcomer reactor was not subjected to gas-solid separation but was directly introduced into the fluidized reaction section of the fluidized bed reactor, so that the first mixed product, the second mixed product and the regenerated B-type catalyst in the fluidized reaction section were brought into contact with each other to carry out a catalytic cracking reaction; and the hydrogenated reformed diesel was no longer introduced into the riser reactor, and the raw material in the riser reactor was only olefin-rich light gasoline.

[0109] Table 1

[0110]

[0111]

[0112] Table 2

[0113]

[0114]

[0115] Table 3 Reaction operating conditions of the embodiments and comparative examples

[0116]

[0117]

[0118] Table 4 is the reaction results of the examples and comparative examples

[0119]

[0120]

[0121] As shown in Table 3 and Table 4, harsh reaction conditions are used in Comparative Example 1 to improve the conversion rate as much as possible, but the coke yield is very high. In Example 1, the down-pipe reactor and the riser reactor are used to effectively control the conversion rate of heavy oil, while maximizing the conversion rate of gasoline and hydro-modified diesel, and combined with the use of the hydro-modified reactor, the diesel yield is reduced, indicating that more diesel is converted into light olefins, and the diesel cetane number is improved, indicating that the diesel quality is improved. Although the light olefin yield is reduced, the slurry yield is higher (i.e., the yield of marine fuel oil is significantly improved), and the coke yield is significantly reduced, which improves the ratio of light olefin yield to coke yield. Compared with Comparative Example 1, the ratio of the sum of ethylene and propylene yields to coke yield in Example 1 is increased from 1.43 to 2.72, which is significantly increased. Compared with Example 1, Example 2 adds hydrogenated reformed diesel and gasoline entering the riser reactor, the yields of ethylene and propylene are further increased, the ratio of the sum of the yields of ethylene and propylene to the coke yield remains substantially unchanged, the cetane number of diesel is increased, the quality of diesel is further improved, and the yields of diesel and gasoline are further reduced. The method of the present invention can produce more marine fuel oil and reduce the coke yield, ensure the yield of light olefins, and increase the ratio of light olefin yield to coke yield.

[0122] The preferred embodiments of the present invention are described in detail above in conjunction with the accompanying drawings. However, the present invention is not limited to the specific details in the above embodiments. Within the technical concept of the present invention, a variety of simple modifications can be made to the technical solution of the present invention, and these simple modifications all belong to the protection scope of the present invention.

[0123] It should also be noted that the various specific technical features described in the above specific embodiments can be combined in any suitable manner without contradiction. In order to avoid unnecessary repetition, the present invention will not further describe various possible combinations.

[0124] In addition, various embodiments of the present invention may be arbitrarily combined, and as long as they do not violate the concept of the present invention, they should also be regarded as the contents disclosed by the present invention.

Claims

1. A method for producing light olefins and marine fuel oil by catalytic cracking, the method comprising: S1, feeding the heavy raw material into the upper part of the descending tube reactor (2) to contact with the first catalytic cracking catalyst from the top of the descending tube reactor (2) and performing a first catalytic cracking reaction from top to bottom to obtain a first mixed product; the conditions of the first catalytic cracking reaction include: a temperature of 420-600° C., a catalyst-oil ratio of 2-14, and a reaction time of 0.2-4 seconds; S2, sending the first mixed product into the settling section (4) of the fluidized bed reactor (3) for gas-solid separation to obtain first oil and gas and a first semi-spent catalyst, and drawing out the first oil and gas from the top of the settling section (4); S3, feeding the light raw material into the lower part of the riser reactor (9) to contact with the second catalytic cracking catalyst and perform a second catalytic cracking reaction from bottom to top to obtain a second mixed product; S4, sending the second mixed product into the fluidized reaction section of the fluidized bed reactor (3) to contact the first semi-regenerated catalyst and the third catalytic cracking catalyst and perform a third catalytic cracking reaction to obtain a third mixed product; separating the third mixed product into a third oil gas and a regenerated catalyst in the settling section (4); drawing out the third oil gas from the top of the settling section (4); sending the regenerated catalyst into a catalyst regenerator (7) for regeneration to obtain a regenerated catalyst; S5, sending the first oil gas and the third oil gas from the settling section (4) to a separation device (12) for separation, sending the separated diesel to a hydro-reforming reactor (13) to contact with a hydrogenation catalyst for hydro-reforming; sending the obtained hydro-reforming diesel to the riser reactor (9) for the second catalytic cracking reaction; The weight ratio of the hydrotreated and reformed diesel fed into the riser reactor to the heavy feedstock fed into the descender reactor is (0.02-0.30):

1.

2. The method according to claim 1, wherein: The conditions of the second catalytic cracking reaction include: temperature of 480-699°C, catalyst-oil ratio of 5-23, and reaction time of 1-10 seconds; The conditions of the third catalytic cracking reaction include: a temperature of 460-640°C, a weight hourly space velocity of 1-35 hours -1 , the reaction pressure is 0.15-0.35 MPa; The conditions of the hydrogenation reforming reaction include: hydrogen partial pressure of 1.2-8.0 MPa, reaction temperature of 150-300°C, hydrogen-oil volume ratio of 150-300, volume space velocity of 1-20 h -1 .

3. The method according to claim 1, wherein: In step S2, the first mixed product is sent to the middle part of the settling section (4) of the fluidized bed reactor (3) for gas-solid separation.

4. The method according to claim 1, wherein: The weight ratio of the light raw material fed into the riser reactor to the heavy raw material fed into the descender reactor is (0.05-0.30): 1; The weight ratio of the hydrotreated and reformed diesel fed into the riser reactor to the heavy feedstock fed into the descender reactor is (0.05-0.20):

1.

5. The method according to claim 1, wherein: The weight ratio of the light raw material fed into the riser reactor to the heavy raw material fed into the downcomer reactor is (0.08-0.20):

1.

6. The method according to claim 1, wherein: Step S4 also includes: sending part of the regenerated catalyst from the catalyst regenerator (7) as the first catalytic cracking catalyst to the top of the downcomer reactor (2); and sending the remaining part of the regenerated catalyst as the second catalytic cracking catalyst and the third catalytic cracking catalyst to the fluidized reaction section of the riser reactor (9) and the fluidized bed reactor (3), respectively.

7. The method according to claim 6, wherein: Based on the total weight of the regenerated catalyst leaving the catalyst regenerator (7) per unit time, 10-70% by weight of the regenerated catalyst is fed into the downcomer reactor (2), 8-30% by weight of the regenerated catalyst is fed into the fluidized reaction section of the fluidized bed reactor (3), and 10-40% by weight of the regenerated catalyst is fed into the riser reactor (9); the sum of the proportions of the regenerated catalyst fed into each reactor is 100% by weight.

8. The method according to claim 1, wherein: Step S4 also includes: first sending the catalyst to be regenerated into the stripping section (5) of the fluidized bed reactor (3) for stripping, and then sending the stripped catalyst to be regenerated into the catalyst regenerator (7) for regeneration.

9. The method according to claim 1, wherein: The regenerated catalyst contains molecular sieve, inorganic oxide and clay; based on the dry weight of the regenerated catalyst, the content of the molecular sieve is 2-51% by weight, the content of the inorganic oxide is 6-97% by weight, and the content of the clay is 5-75% by weight; the sum of the contents of the molecular sieve, the inorganic oxide and the clay is 100% by weight; The molecular sieves include a shape-selective molecular sieve with an average pore size of less than 0.7 nanometers and a Y-type molecular sieve; on a dry basis and based on the total weight of the molecular sieves, the content of the shape-selective molecular sieve with an average pore size of less than 0.7 nanometers is 45-95% by weight, and the content of the Y-type molecular sieve is 2-54% by weight; The shape-selective molecular sieve with an average pore size of less than 0.7 nanometers is selected from one or more of the ZSM series molecular sieves, ZRP molecular sieves, ferrierite, chabazite, cyclopentane, erionite, A zeolite, column zeolite and lanoite, and the Y-type molecular sieve is selected from one or more of the rare earth Y-type molecular sieve, rare earth hydrogen Y-type molecular sieve, ultra-stable Y-type molecular sieve and rare earth ultra-stable Y-type molecular sieve.

10. The method according to claim 1, wherein: The hydrogenation catalyst contains a carrier and an active metal component; the content of the active metal component is 1-20% by weight based on the dry weight of the hydrogenation catalyst; The carrier is selected from alumina and / or amorphous silica-alumina, and the active metal component is selected from non-noble metals of Group VIB and / or non-noble metals of Group VIII.

11. The method according to claim 1, wherein: In step S1, the heavy raw material is one or more selected from vacuum wax oil, atmospheric residue oil, coker wax oil, deasphalted oil, furfural refined raffinate oil, coal liquefaction oil, oil sand oil and shale oil; In step S3, the light feedstock is an olefin-rich liquefied gas component and / or an olefin-rich light gasoline component; the hydrotreated and reformed diesel fed into the riser reactor for the second catalytic cracking reaction is part or all of the heavy fraction from the hydrotreatment and reforming reactor.

12. A system for the method for producing light olefins and marine fuel oil according to any one of claims 1 to 11, the system comprising a downcomer reactor (2), a fluidized bed reactor (3), a catalyst regenerator (7), a riser reactor (9), a separation device (12) and a hydro-upgrading reactor (13); The downtube reactor (2) is provided with a catalyst inlet at the top, a heavy feedstock inlet at the upper part and a product outlet at the bottom; the fluidized bed reactor (3) comprises a settling section (4) and a fluidized reaction section which are fluidically connected, the fluidized reaction section being arranged below the settling section (4), and a first material inlet arranged at the settling section, a gas outlet arranged above the settling section, a second material inlet arranged at the fluidized reaction section, a catalyst inlet and a catalyst outlet arranged at the fluidized reaction section; the catalyst regenerator (7) is provided with a catalyst inlet and a catalyst outlet; the riser reactor (9) is provided with a catalyst inlet at the bottom, a feedstock inlet at the lower part and a product outlet at the top; the separation device (12) is provided with a material inlet and a diesel outlet, and the hydro-reforming reactor (13) is provided with a material inlet and a product outlet; The product outlet of the downtube reactor (2) is connected to the first material inlet of the fluidized bed reactor (3), the product outlet of the riser reactor (9) is connected to the second material inlet of the fluidized bed reactor (3), the catalyst inlet of the catalyst regenerator (7) is connected to the catalyst outlet of the fluidized bed reactor (3), the gas outlet of the fluidized bed reactor (3) is connected to the inlet of the separation device (12), the diesel outlet of the separation device (12) is connected to the material inlet of the hydro-modification reactor (13), and the product outlet of the hydro-modification reactor (13) is connected to the raw material inlet of the riser reactor (9).

13. The system according to claim 12, wherein: The catalyst inlet of the fluidized bed reactor (3), the catalyst inlet of the downcomer reactor (2) and the catalyst inlet of the riser reactor (9) are each independently connected to the catalyst outlet of the catalyst regenerator (7).

14. The system according to claim 12, wherein: The fluidized bed reactor (3) further comprises a stripping section (5) in fluid communication with the fluidized reaction section. The stripping section (5) is arranged downstream of the fluidized reaction section. The catalyst outlet of the fluidized bed reactor (3) is arranged at the lower part of the stripping section (5).

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