CN116925810A

CN116925810A - Catalytic cracking method and catalytic cracking device for raw oil

Info

Publication number: CN116925810A
Application number: CN202210351484.9A
Authority: CN
Inventors: 吴青
Original assignee: China National Offshore Oil Corp CNOOC
Current assignee: China National Offshore Oil Corp CNOOC
Priority date: 2022-04-02
Filing date: 2022-04-02
Publication date: 2023-10-24

Abstract

The invention relates to the technical field of catalytic cracking, and discloses a method for catalytic cracking raw oil and a catalytic cracking device, wherein the method comprises the following steps: 1) The raw oil contacts with DPC1-a catalyst, flows from top to bottom in a first reaction zone, and undergoes a first reaction to obtain a first mixture; 2) The first mixture is contacted with DPC2-a catalyst and recycle oil, flows from bottom to top in a second reaction zone, and undergoes a second reaction to obtain a second mixture; 3) Separating the second mixture to obtain a spent mixed catalyst and reaction oil gas; 4) Carrying out a first regeneration reaction on the mixed catalyst to be regenerated to obtain a regenerated DPC2-a catalyst and a semi-regenerated DPC1-a catalyst; 5) Carrying out a second regeneration reaction on the semi-regenerated DPC1-a catalyst to obtain a regenerated DPC1-a catalyst; 6) Fractionating the reacted oil gas to obtain dry gas, liquefied gas, gasoline, diesel oil and heavy fraction. The method widens the source of catalytic cracking raw oil, and can maximize production of C ₂ ‑C ₄ Olefins, and clean gasoline is obtained.

Description

Catalytic cracking method and catalytic cracking device for raw oil

Technical Field

The invention relates to the technical field of catalytic cracking, in particular to a method for catalytic cracking of raw oil and a catalytic cracking device.

Background

Catalytic cracking is a core device for converting heavy oil into light oil and low-carbon olefin in oil refineries, and catalytic cracking has become an important means for obtaining economic benefit for oil refineries through development of nearly a century. With the continuous promotion of the increasing weight of crude oil and environmental protection regulations in the world and the requirement of improving the oil refining benefit, catalytic cracking has put higher demands on improving the conversion of inferior raw oil, the yield of target products, the quality of gasoline and diesel oil, the production of high-yield low-carbon olefin chemicals and the like.

Early catalytic cracking uses vacuum distillate as a main raw material, and can also be mixed with coking heavy distillate, but the vacuum distillate of some crude oil has high sulfur content, little hydrogen content or low characteristic factor, and is not suitable for being directly used as a catalytic cracking raw material. Along with the gradual heavy and poor quality of crude oil in the world, refineries in various countries develop the fluid catalytic cracking technology for blending or fully blending heavy oil and residual oil so as to widen the source of FCC raw oil, meet the demand of the global market for light chemicals and furthest improve the economic benefit.

Around the poor quality heavy oil conversion and oil refining chemical integration, various large oil refineries and research institutes successively develop a plurality of new catalytic cracking processes.

MSCC technology of UOP: MSCC is a downward reaction technology, and the catalyst contacts with the ascending oil gas from top to bottom on a plane and then is separated rapidly. The system can weaken the slipping phenomenon of the catalyst as much as possible, reduce the secondary cracking and thermal cracking reactions, improve the catalyst-to-oil ratio, reduce the regeneration temperature and omit a catalyst heater. The MSCC process oil has short contact time, improved dry gas selectivity, improved conversion rate of raw oil, obviously improved gasoline selectivity, increased yields of liquefied gas, gasoline and diesel oil, and enhanced tolerance of catalyst to Ni. But the MSCC technology has poor adaptability to high carbon residue, high boiling point or high slag mixing ratio raw materials, and is suitable for light clean wax oil.

SCT technology of ExxonMobil: the outlet area of the process reactor adopts a closed coupling cyclone separator of Exxon company patent technology, so that the catalyst and the cracked products can be rapidly separated, and the operation stability is good; the stripper adopts an advanced sectional stripping device, so that hydrocarbons carried on the catalyst can be removed better, and coke generation is reduced; the feeding system improves atomization performance and agent-oil contact condition, and reduces back mixing.

The petrocat process of Petrobras: the process uses a catalyst cooler, one part of the cooled regenerated catalyst returns to the bed layer of the regenerator, the other part is mixed with the hot regenerated catalyst, and the temperature of the regenerated catalyst after mixing is controlled, so that the mixed regenerated catalyst with two paths of different temperatures contacts with the raw oil. The temperature after mixing is controlled to be about 650 ℃, and higher conversion rate can be obtained when the temperature of the raw materials is higher. When the catalyst and the raw materials are mixed, the temperature difference between the regenerated catalyst and the feed is reduced from 400-450 ℃ to 290 ℃, the reaction proportion in a high-temperature zone is reduced, and the tendency of thermal cracking reaction is greatly reduced. On the other hand, more heat comes from the preheating of the feedstock oil, shortening the hydrocarbon gasification time, and thus reducing the thermal reaction.

HS-FCC process: the process adopts a downlink reactor, and the plug flow has no back mixing and concentrated residence time distribution, thereby being beneficial to reaction selectivity, eliminating pipe wall channeling and reducing thermal reaction. The reaction contact time of the technology is lower than 0.5 seconds, so that the thermal reaction can be reduced again, the serial secondary reaction can be inhibited, and the coking can be inhibited; the HS-FCC process agent has high oil ratio, and can further reduce thermal reaction, promote conversion and inhibit side reaction.

Dan Ke, MIP technology: the process mainly aims at reducing the olefin content of gasoline, the reactor is divided into two reaction areas, the first reaction area is mainly used for primary cracking reaction, higher reaction temperature and catalyst-to-oil ratio are adopted to generate more olefin and process heavier raw oil, the raw oil enters a second expanded reaction area after shorter residence time, the second reaction area reduces the flow rate of oil gas and catalyst, inhibits secondary cracking reaction, increases isomerization and hydrogen transfer reaction, increases the isoparaffin and arene content in the gasoline, and reduces the olefin content; the process can reduce the olefin content of the gasoline to 18-24%, meets the national V gasoline standard, but cannot meet the national VI gasoline standard.

Dan Ke institute DCC process: the process uses heavy oil as raw material, uses solid acid shape-selective molecular sieve catalyst, and uses high yield of ethylene and propylene as main purpose. The DCC-I selects more severe operating conditions, and reacts in a riser encrypted phase fluidized bed reactor to produce the gas olefin with propylene as the main component in maximum; the DCC-II selects mild operation conditions, reacts in a riser reactor to produce small molecular olefins such as propylene, isobutene, isoamylene and the like to the maximum extent, and simultaneously produces high-octane and high-quality gasoline.

However, in the conventional catalytic cracking units, acidic molecular sieve catalysts are adopted, and if poor-quality heavy raw materials containing heavy metals such as nickel, vanadium, calcium, iron and the like and coked wax oil containing high alkaline nitrogen are directly used as raw materials, the yield of target products (low-carbon olefins, gasoline and diesel oil) is seriously reduced, and the catalyst consumption is greatly increased. Therefore, heavy metals and basic nitrogen must be removed by a pretreatment device such as heavy oil hydrogenation and the like before the heavy metals and basic nitrogen can be used as raw materials of a catalytic device,

pretreatment of inferior heavy oil is a widely used means for widening catalytic raw material sources at present. The hydrogenation treatment method is adopted at first, the desulfurization is mainly adopted, and the conversion rate of VGO can reach 25% -40% under the medium pressure (5-8 MPa) with the improvement of the catalyst and the development of the hydrocracking process, so that the desulfurization index is met, and the product quality is improved greatly. In the last 80 th century, the oil refining industry began to pay attention to the conversion of residual oil fractions with boiling points higher than 500-550 ℃ in heavy oil, and besides the traditional thermal conversion-coking process, various conversion processes and separation processes were developed, and the process technologies of Solvent Deasphalting (SDA), mild and moderate pyrolysis of vacuum residual oil, deep pyrolysis, hydrodesulfurization of residual oil, hydrocracking and the like were used for removing asphaltenes and heavy colloid (too high sulfur, nitrogen and heavy metals, low H/C ratio and high carbon residue value) in residual oil, which could not be used as a catalytic cracking raw material, to obtain qualified catalytic cracking raw oil with higher yield. In addition, the LC-fining process, the Hycon process, the suspension bed hydrocracking process such as VCC process, SOC process, canmet process, etc. can be used for preparing the catalytic cracking raw material, but the hydrocracking process has complex process flow, large hydrogen consumption, more equipment, high one-time investment and high raw material pretreatment cost.

Therefore, there is a need to provide a method for obtaining low-olefin clean gasoline meeting the national VI standard, which can directly perform catalytic cracking without pretreatment of heavy and inferior raw oil with high impurity content.

Disclosure of Invention

In order to achieve the above object, a first aspect of the present invention provides a catalytic cracking method of a raw oil, the method comprising the steps of:

(1) The raw oil contacts with DPC1-a catalyst, and flows from top to bottom in a first reaction zone under the carrying of a fluidization medium to perform a first reaction to obtain a first mixture;

(2) The first mixture is carried by a pre-lifting medium, contacts with a DPC2-a catalyst and recycle oil, flows from bottom to top in a second reaction zone, and undergoes a second reaction to obtain a second mixture; wherein the DPC1-a catalyst has a stronger basicity than the DPC2-a catalyst, and the DPC1-a catalyst has a lower bulk ratio than the DPC2-a catalyst;

(3) Separating the second mixture to obtain a spent mixed catalyst and reaction oil gas;

(4) Carrying out a first regeneration reaction after the mixed catalyst to be regenerated is subjected to preliminary removal of entrained reaction oil gas to obtain a regenerated DPC2-a catalyst and a semi-regenerated DPC1-a catalyst; wherein the regenerated DPC2-a catalyst obtained is recycled back to step (2);

(5) Carrying out a second regeneration reaction on the semi-regenerated DPC1-a catalyst to obtain a regenerated DPC1-a catalyst; wherein the regenerated DPC1-a catalyst obtained is recycled to step (1);

(6) Fractionating the reaction oil gas to obtain dry gas, liquefied gas, gasoline, diesel oil and heavy fraction; wherein the heavy oil fraction is returned to the second reaction zone as recycle oil for recycle.

In a second aspect, the invention provides a catalytic cracking device, wherein the device comprises a downer reaction tube (1), a riser (2), a settler (3), a first regenerator (4) and a second regenerator (5) which are connected end to end in sequence.

Through the technical scheme, the beneficial technical effects obtained by the invention are as follows:

1) The catalytic cracking method of the raw oil breaks through the technical bottleneck of the traditional catalytic cracking device using an acid catalyst, has good adaptability to raw materials, can directly treat poor heavy oil with heavy metals such as high nickel content, vanadium content, iron content, calcium content and the like, can directly treat coker gas oil with high sulfur content and nitrogen content, can directly treat crude oil with high pollution impurities content, and widens the source of the catalytic cracking raw oil;

2) According to the catalytic cracking method of the raw oil, disclosed by the invention, the raw material does not need to be subjected to hydrofining pretreatment, so that the cost of producing low-carbon olefin and clean gasoline in an oil refinery can be greatly reduced;

3) The catalytic cracking method of the raw oil provided by the invention realizes the maximum production of C by heavy raw materials ₄ The following low-carbon olefins (dry gas and liquefied gas) can be obtained at the same time, clean gasoline with low olefin content can be obtained.

Drawings

Fig. 1 is a schematic view of a catalytic cracker in a preferred embodiment of the present invention.

Description of the reference numerals

1, a downer reaction tube 2, a riser 3 and a settler

4, a first regenerator 5, a second regenerator

Detailed Description

The endpoints and any values of the ranges disclosed herein are not limited to the precise range or value, and are understood to encompass values approaching those ranges or values. For numerical ranges, one or more new numerical ranges may be found between the endpoints of each range, between the endpoint of each range and the individual point value, and between the individual point value, in combination with each other, and are to be considered as specifically disclosed herein.

The first aspect of the invention provides a catalytic cracking method of raw oil, wherein the method comprises the following steps:

In step (1):

in the invention, the first reaction of the raw oil contacted with the DPC1-a catalyst comprises a cracking shearing reaction and a refining reaction, and is used for removing substances such as sulfur, nitrogen, metal, carbon residue and the like in the raw oil so as to improve the catalytic cracking effect of the subsequent raw oil, improve the distribution of chemicals and improve the yield and quality of the chemicals. The inventor of the invention finds that by changing the flow direction of materials in the reaction process and coupling a first reaction zone flowing from top to bottom with a second reaction zone flowing from bottom to top, the C in a first reaction product can be promoted ₅ The olefins continue to undergo reactions such as cracking, aromatization, cyclization, hydrogen transfer and the like to generate C ₄ The following small molecular olefins can directly obtain dry gas rich in ethylene and C ₃ Olefins and C ₄ Liquefied gas of olefin, clean gasoline of low olefin, diesel oil, heavy fraction, etc.

In a preferred embodiment, the feedstock is selected from crude oil and/or heavy oil; the crude oil refers to petroleum which is not extracted, the heavy oil refers to residues left after the crude oil is processed and light components are extracted, and the residues can be selected from various types of intermediate, intermediate-cycloalkyl and cycloalkyl heavy distillate oil such as straight-run wax oil, coker wax oil, hydrogenated wax oil, atmospheric residue, vacuum residue and the like.

In the method for directly preparing chemicals from the raw oil, provided by the invention, the raw oil has strong adaptability, and crude oil is not required to be specially limited. Specifically, when the base property of crude oil is classified, the crude oil may be classified into paraffinic crude oil, intermediate base crude oil, intermediate-naphthenic crude oil and naphthenic crude oil, and the method for directly preparing chemicals from raw oil provided by the present invention may be used for treating both paraffinic crude oil and intermediate base crude oil, intermediate-naphthenic crude oil and naphthenic crude oil. When the density of crude oil is divided, the crude oil can be divided into light crude oil, medium crude oil, heavy crude oil and extra thick crude oil. The raw oil in the invention can be single crude oil or single heavy oil, or multiple crude oils, multiple heavy oils or mixed oil of crude oil and heavy oil. That is, the raw oil is one or more selected from oil sand asphalt, venezuela extra heavy oil, straight run wax oil, coker wax oil, hydrogenated wax oil, atmospheric residue and vacuum residue.

In a preferred embodiment, the DPC1-a catalyst comprises 85-99 parts by weight of carrier I and 1-15 parts by weight of active metal oxide I; wherein the carrier I is at least one selected from alumina, silica, titania and zirconia; the active metal oxide I is selected from alkali metal oxide and/or alkaline earth metal oxide.

In a preferred embodiment, the DPC1-a catalyst comprises 90-98 parts by weight, preferably 94-97 parts by weight, of carrier I and 2-10 parts by weight, preferably 3-6 parts by weight, of active metal oxide I.

In a preferred embodiment, the support I is selected from alumina and/or silica, preferably silica.

In a preferred embodiment, the active metal oxide I is selected from at least one of calcium oxide, magnesium oxide, barium oxide, strontium oxide, preferably magnesium oxide and/or barium oxide.

In a preferred embodimentIn embodiments, the DPC1-a catalyst is CO ₂ The desorption peak temperature is 185-195 ℃, preferably 187-192 ℃; CO ₂ The number of base centers at the desorption peak position is 16 to 22mmol/g, preferably 18 to 21mmol/g.

Wherein, in the present invention, the CO of the DPC1-a catalyst ₂ Peak desorption temperature and CO ₂ The number of base centers at the desorption peak position was determined by the temperature-programmed carbon dioxide adsorption method (CO ₂ TPD) was tested on a Quantachrome ChemBet 3000 chemisorber. Taking 150mg of catalyst sample, pretreating at 600 ℃ for 1h under He gas atmosphere, and then cooling to 100 ℃ for CO ₂ And (5) adsorption. The volume ratio of the use is 1: CO of 9 ₂ The mixture of He and the catalyst is used as adsorption gas, adsorbed for 30min at 100 ℃, and then purged for 30min by He gas to remove physically adsorbed CO ₂ . Finally, under the He atmosphere, the temperature is 16 ℃ for min ^-1 Is desorbed at a rate of from 100 ℃ to 600 ℃ to obtain CO ₂ -TPD profile. From CO ₂ Reading CO in TPD profile ₂ The desorption peak temperature, calculate CO ₂ Number of base centers at the desorption peak position. Wherein, in the present invention, CO ₂ The higher the desorption peak temperature, the higher the CO ₂ The more the number of base centers at the desorption peak position, the more basic the catalyst.

In a preferred embodiment, the DPC1-a catalyst has a bulk ratio of 0.5 to 0.7g/mL, preferably 0.55 to 0.65g/mL; the particle size is 30-100. Mu.m, preferably 50-70. Mu.m.

In a preferred embodiment, the fluidizing medium is selected from at least one of steam, dry gas, natural gas and liquefied gas, preferably steam.

In a preferred embodiment, the feed temperature of the feedstock is 180-340 ℃, preferably 200-330 ℃; the feed temperature of the DPC1-a catalyst is 580-750 ℃, preferably 600-670 ℃; the feed temperature of the fluidization medium is 200-330 ℃, preferably 250-300 ℃.

In a preferred embodiment, the mass ratio of the DPC1-a catalyst to the feedstock oil is from 5 to 20:1, preferably 8-15:1, a step of; the mass ratio of the raw oil to the fluidization medium is 100:1-15, preferably 100:1-8.

In a preferred embodiment, the temperature of the first reaction is 370-535 ℃, preferably 410-525 ℃; the pressure of the first reaction is 0.1-1MPa, preferably 0.1-0.4MPa; the time of the first reaction is 0.1 to 4s, preferably 0.5 to 3s.

In step (2):

wherein in the present invention, the first reaction occurring in the first reaction zone flowing from top to bottom is the second reaction occurring in the second reaction zone flowing from bottom to top, further reduces C ₅ The olefin product has increased C ₄ The following olefin products create conditions, and the first mixture can promote the complex reactions of cracking, isomerization, aromatization, disproportionation, cyclization, hydrogen transfer, transalkylation, condensation and the like in the second reaction zone under the combined action of the DPC1-a catalyst and the DPC2-a catalyst, so as to lighten heavy components and obtain dry gas and liquefied gas rich in low-carbon olefin, low-olefin clean gasoline, diesel oil and heavy fractions.

In a preferred embodiment, the first mixture comprises all of the material after the reaction in the first reaction zone; the pre-lifting medium is selected from at least one of steam, dry gas, natural gas and liquefied gas, preferably steam.

In a preferred embodiment, the DPC2-a catalyst preferably comprises 54-63 parts by weight of carrier II,1-5 parts by weight of active metal oxide II and 33-43 parts by weight of molecular sieve; wherein the carrier II is at least one of alumina, silica, titania and zirconia; the active metal oxide II is selected from alkali metal oxide and/or alkaline earth metal oxide; the molecular sieve comprises a medium pore molecular sieve and a large pore molecular sieve, wherein the medium pore molecular sieve is selected from ZSM-5 molecular sieve and/or ZRP-1 molecular sieve, and the large pore molecular sieve is selected from REHY type molecular sieve and/or REUSY type molecular sieve.

In a preferred embodiment, the DPC2-a catalyst preferably comprises 55-62 parts by weight of carrier II,1.5-2.5 parts by weight of active metal oxide II and 35-40 parts by weight of molecular sieve.

In a preferred embodiment, the support II is selected from alumina and/or silica, preferably silica.

In a preferred embodiment, the active metal oxide II is selected from at least one of calcium oxide, magnesium oxide, barium oxide, strontium oxide, preferably calcium oxide and/or magnesium oxide.

In a preferred embodiment, the support I and the metal active component I in the DPC1-a catalyst are the same as the support II and the metal active component II in the DPC2-a catalyst.

In a preferred embodiment, the mass ratio of the medium pore molecular sieve to the large pore molecular sieve is 100:3-15, preferably 100:5-12.

The silicon-aluminum ratio of ZSM-5 molecular sieve, ZRP-1 molecular sieve, REHY molecular sieve and REUSY molecular sieve is not particularly limited, and the silicon-aluminum ratio commonly used in the molecular sieves in the field can be used in the invention. The sources of the ZSM-5 molecular sieve, the ZRP-1 molecular sieve, the REHY molecular sieve and the REUSY molecular sieve are not particularly limited, and the molecular sieve can be a commercial product or can be prepared according to a conventional method in the field. The preparation method of the molecular sieve is not repeated in the invention.

In a preferred embodiment, the ZSM-5 molecular sieve has a silica/alumina mole ratio of from 20 to 100:1, preferably 25-50:1, a step of; the mole ratio of silicon oxide/aluminum oxide in the ZRP-1 molecular sieve is 40-150:1, preferably 50-120:1, a step of; the mol ratio of silicon oxide/aluminum oxide in the REHY type molecular sieve and the REUSY type molecular sieve is 1-30:1, preferably 3-15:1.

In a preferred embodiment, the mesoporous molecular sieve is selected from the group consisting of ZSM-5 molecular sieves and ZRP-1 molecular sieves; wherein the mass ratio of the ZSM-5 molecular sieve to the ZRP-1 molecular sieve is 8-20:1, preferably 10-15:1.

in a preferred embodiment, the large pore molecular sieve is a REUSY type molecular sieve.

In a preferred embodiment, the DPC2-a catalyst is CO ₂ The desorption peak temperature is 160-190 ℃, preferably 170-185 ℃; CO ₂ Number of base centers at the desorption peak position5-18mmol/g, preferably 7-15mmol/g.

In a preferred embodiment, the DPC2-a catalyst has a bulk ratio of 0.65-0.95g/mL, preferably 0.75-0.85g/mL; the particle size is 40-120. Mu.m, preferably 70-90. Mu.m.

In a preferred embodiment, the DPC1-a catalyst is CO ₂ The number of base centers at the desorption peak position is at least 10mmol/g more than that of the DPC2-a catalyst; the DPC2-a catalyst has a bulk ratio at least 0.15-0.25g/mL greater than the DPC1-a catalyst.

In a preferred embodiment, the cycle oil is a heavy fraction derived from the reaction oil and gas.

In a preferred embodiment, the recycle oil is fed at a temperature of 200 to 350 ℃, preferably 250 to 330 ℃; the feed temperature of the DPC2-a catalyst is 660-750 ℃, preferably 680-720 ℃; the feed temperature of the pre-lifting medium is 200-300 ℃, preferably 230-280 ℃.

In a preferred embodiment, the mass ratio of recycle oil to feed oil (recycle ratio) is from 0.1 to 0.5, preferably from 0.2 to 0.4; the mass ratio of the raw oil to the pre-lifting medium is 100:1-15, preferably 100:1-8.

In a preferred embodiment, the reaction conditions of the second reaction further comprise: the DPC2-a catalyst is used in an amount such that the reaction temperature of the second reaction is 500-560 ℃, preferably 515-530 ℃; the pressure of the second reaction is 0.1-0.5MPa, preferably 0.2-0.3MPa; the time for the second reaction is 0.5 to 5s, preferably 1 to 3s.

Wherein, on the premise of determining the feeding quantity of the first mixture, the recycle oil and the pre-lifting medium and the feeding temperature, the invention controls the reaction temperature of the second reaction by controlling the feeding quantity of the DPC2-a catalyst with the feeding temperature determined.

In step (3):

in a preferred embodiment, the separation is performed in a settler. The separation is not particularly limited, and may be performed according to a conventional operation in the art.

In a preferred embodiment, the second mixture is separated by contacting with stripping steam, wherein the mass ratio of feedstock oil to stripping steam is 100:1-20, preferably 100:1-10.

In step (4):

in a preferred embodiment, the first regeneration reaction comprises a first calcination under an air atmosphere; preferably, the temperature of the first firing is 650-730 ℃, preferably 680-720 ℃; the pressure of the first roasting is 0.1-0.5MPa, preferably 0.15-0.35MPa; the dense bed flue gas line velocity of the first regenerator is 0.3-1m/s, preferably 0.5-0.9m/s; the linear velocity of the flue gas of the dilute phase bed in the first regenerator is 0.8-1.8m/s, preferably 1.1-1.5m/s.

In step (5):

in a preferred embodiment, the second regeneration reaction comprises a second calcination under an atmosphere of air and/or water vapor; preferably, the temperature of the second firing is 650-720 ℃, preferably 660-700 ℃; the pressure of the second roasting is 0.1-0.5MPa, preferably 0.15-0.35MPa; the dense bed flue gas linear velocity of the second regenerator is 0.3-0.9m/s, preferably 0.5-0.8m/s; the flue gas linear velocity of the dilute phase bed of the second regenerator is 0.4-1.1m/s, preferably 0.5-0.9m/s.

In step (6):

in a preferred embodiment, the fractionation is not particularly limited, and may be performed according to a conventional operation in the art, and the present invention will not be described in detail. Preferably, the distillation range of the gasoline fraction is 30-205 ℃, the distillation range of the diesel fraction is 206-380 ℃, and the heavy fraction is a distillate oil with the temperature above 380 ℃.

In a second aspect the invention provides a catalytic cracking unit comprising a downer reaction tube 1, a riser 2, a settler 3, a first regenerator 4 and a second regenerator 5 connected in series, as particularly shown in figure 1.

In a preferred embodiment of the present invention, the second regenerator 5 is connected to the downer reaction tube 1 via a DPC1-a catalyst feed.

In a preferred embodiment of the present invention, the downer reaction tube 1 is provided with a raw oil inlet for introducing raw oil; wherein the raw oil inlet is arranged below the joint of the DPC1-a catalyst feeding pipe and the downer reaction pipe.

In a preferred embodiment of the present invention, a fluidization medium inlet is further provided on the downer reaction tube 1, wherein the fluidization medium inlet is provided above the connection of the DPC1-a catalyst feed tube and the downer reaction tube.

The fluidized medium inlet is used for introducing a fluidized medium, the fluidized medium carries the DPC1-a catalyst and raw oil to flow in the riser reaction tube from top to bottom, and a first reaction occurs to obtain a first mixture.

In a preferred embodiment of the present invention, the lower end of the downer reaction tube 1 is connected to the lower end of the riser 2, and the downer reaction tube 1 and the riser 2 are in a "U" shape. In the invention, the downer reaction tube comprises a vertical tube and an inclined tube, wherein the inclined tube is used for connecting the vertical tube and the lifting tube, so that the downer reaction tube and the lifting tube are U-shaped.

In a preferred embodiment, a DPC2-a catalyst inlet and a recycle oil inlet are provided at the lower end of the riser 2; further preferably, the DPC2-a catalyst inlet and cycle oil inlet are each independently disposed above the junction of the downer reaction tube 1 and the riser 2. Preferably, the DPC2-a catalyst inlet is connected to the first regenerator 4 through a DPC2-a catalyst feed line.

In a preferred embodiment, the lower end of the riser 2 is further provided with a pre-lift medium inlet; wherein the pre-lifting medium inlet is arranged below the joint of the downer reaction tube 1 and the lifting tube 2.

In the invention, the first mixture from the downer reaction tube 1 directly enters the riser tube 2 without separation operation, contacts with the DPC2-a catalyst and recycle oil, flows from bottom to top in the riser tube 2 under the action of a pre-lifting medium, and undergoes a second reaction to obtain a second mixture.

In a preferred embodiment, the upper end of the riser 2 is connected to a settler 3. In the invention, the obtained second mixture directly enters a settler for separation to obtain the spent mixed catalyst and the reaction oil gas. Wherein the spent mixed catalyst comprises a spent DPC1-a catalyst and a spent DPC2-a catalyst.

In a preferred embodiment, the settler 3 is connected to the first regenerator 4 via a spent mixed catalyst discharge pipe, the first regenerator (4) is connected to the riser DPC2-a catalyst inlet via a DPC2-a catalyst feed pipe, and the first regenerator 4 is connected to the second regenerator 5 via a semi-regenerated DPC1-a catalyst feed pipe.

Wherein in the present invention, the DPC1-a catalyst has a smaller bulk ratio than the DPC2-a catalyst, and the first regenerator is used for regenerating the to-be-regenerated DPC2-a catalyst and separating out the semi-regenerated DPC1-a catalyst. The regenerated DPC2-a catalyst obtained after regeneration is returned to the riser for recycling. The separated semi-regenerated DPC1-a catalyst enters a second regenerator to be continuously regenerated, and the regenerated DPC1-a catalyst obtained after regeneration returns to a downstream bed reaction tube to be recycled.

In a preferred embodiment, the first regenerator 4 is provided with a circulation pipe for controlling the level stabilization of the dense bed of the first regenerator, ensuring continuous and stable operation of the catalyst regeneration reactor.

In a preferred embodiment, a catalyst inventory adjusting pipe is further provided between the second regenerator 5 and the first regenerator 4, and the catalyst inventory adjusting pipe is used for adjusting the catalyst inventory in the second regenerator.

In a preferred embodiment, the apparatus further comprises a fractionation column connected to the settler, wherein the fractionation column is used to separate the reaction oil and gas to obtain dry gas, liquefied gas, gasoline, diesel and heavy fractions.

The present invention will be described in detail by examples. The example is carried out in a catalytic cracking unit shown in fig. 1, wherein the catalytic cracking unit comprises a downer reaction tube, a riser, a settler, a first regenerator and a second regenerator which are connected end to end in this order. The second regenerator is connected with the downer reaction tube through a DPC1-a catalyst feeding tube, a raw oil inlet is arranged below the connection position of the DPC1-a catalyst feeding tube and the downer reaction tube, and a fluidization medium inlet is arranged above the connection position of the DPC1-a catalyst feeding tube and the downer reaction tube.

The lower end of the downer reaction tube is connected with the lower end of the lifting tube and is U-shaped, a DPC2-a catalyst inlet and a recycle oil inlet are arranged above the joint of the downer reaction tube and the lifting tube, and a pre-lifting medium inlet is arranged below the joint of the downer reaction tube and the lifting tube; the upper end of the lifting pipe is connected with a settler, and the settler is connected with a fractionating tower;

The settler is connected with a first regenerator through a discharging pipe of a mixed catalyst to be regenerated, a circulating pipe is arranged on the first regenerator, the first regenerator is connected with a DPC2-a catalyst inlet on a riser through a DPC2-a catalyst feeding pipe, and the first regenerator is connected with the second regenerator through a semi-regenerated DPC1-a catalyst feeding pipe;

a material level regulating pipe is arranged between the second regenerator and the first regenerator, and the second regenerator is connected with a DPC1-a catalyst inlet on the downer reaction pipe through a DPC1-a catalyst pipe.

The present invention will be described in detail by examples. Wherein the DPC1-a catalyst used in the examples contains 95wt% of silica and 5wt% of magnesia, CO of the DPC1-a catalyst ₂ The desorption peak temperature is 189 ℃, CO ₂ The number of alkali centers at the desorption peak position is 20.27mmol/g, the heap ratio is 0.55g/mL, and the particle size is 60 mu m;

the DPC2-a catalyst comprises 62wt% of silicon oxide, 0.8wt% of calcium oxide, 1.2wt% of magnesium oxide, 30wt% of ZSM-5 molecular sieve (the mol ratio of silicon oxide to aluminum oxide is 30:1), 2.5wt% of ZRP-1 molecular sieve (the mol ratio of silicon oxide to aluminum oxide is 50:1) and 3.5wt% of REUSY type molecular sieve (the mol ratio of silicon oxide to aluminum oxide is 4.8:1); CO of DPC2-a catalyst ₂ The desorption peak temperature is 175 ℃, CO ₂ The number of the alkali centers at the desorption peak position is 9.23mmol/g, and the pile ratio is 0.75g/mL, particle size 80 μm.

The mid-cycloalkyl vacuum residuum and mid-cycloalkyl coker gas oil in the examples were both from the middle sea oil Huiz petrochemical Co.

Example 1

(1) DPC1-a catalyst from the second regenerator enters a downer reaction tube through a DPC1-a catalyst feed tube, raw oil (intermediate-cycloalkyl vacuum residue) enters the downer reaction tube through a raw oil inlet, fluidization medium (water vapor) enters the downer reaction tube through a fluidization medium inlet, the DPC1-a catalyst contacts with the raw oil in the downer reaction tube under the carrying of the fluidization medium, flows from top to bottom in a first reaction zone, and a first reaction occurs to obtain a first mixture; wherein the feed temperature of the raw oil is 250 ℃, the feed temperature of the DPC1-a catalyst is 660 ℃, and the feed temperature of the fluidization medium is 275 ℃; the mass ratio of the DPC1-a catalyst to the raw oil is 8:1, the mass ratio of the raw oil to the fluidization medium is 100:1, a step of; the temperature of the first reaction is 510 ℃, the pressure of the first reaction is 0.26MPa, and the time of the first reaction is 2.0s;

(2) The pre-lifting medium (water vapor) enters a lifting pipe through a pre-lifting medium inlet, the DPC2-a catalyst from the first regenerator enters the lifting pipe through a DPC2-a catalyst feeding pipe, the first mixture is contacted with the DPC2-a catalyst and recycle oil under the carrying of the pre-lifting medium, and flows from bottom to top in a second reaction zone to generate a second reaction to obtain a second mixture; wherein, the recycle oil is heavy fraction, the feeding temperature of the recycle oil is 300 ℃, the feeding temperature of the DPC2-a catalyst is 690 ℃, and the feeding temperature of the pre-lifting medium is 275 ℃; the mass ratio of recycle oil to feed oil (recycle ratio) was 0.3, and the mass ratio of feed oil to pre-lifting medium was 100:2.5; the DPC2-a catalyst was used in such an amount that the temperature of the second reaction was 530℃and the pressure of the second reaction was 0.26MPa, and the time of the second reaction was 3.0s;

(3) The second mixture obtained after the riser reaction enters a settler to be contacted with stripping steam for separation, and a spent mixed catalyst and reaction oil gas are obtained; wherein, the mass ratio of the raw oil to the stripping steam is 100:4.5;

(4) The separated spent mixed catalyst enters a first regenerator through a spent mixed catalyst pipe to carry out a first regeneration reaction, and is subjected to first roasting under the air atmosphere to obtain a regenerated DPC2-a catalyst and a semi-regenerated DPC1-a catalyst; wherein the temperature of the first roasting is 680 ℃, the pressure of the first roasting is 0.27MPa, the dense-phase bed flue gas linear velocity of the first regenerator is 0.6m/s, and the dilute-phase bed flue gas linear velocity is 1.3m/s; the obtained regenerated DPC2-a catalyst is circulated back to the riser through a DPC2-a catalyst feed pipe;

(5) The obtained semi-regenerated DPC1-a catalyst enters a second regenerator through a semi-regenerated DPC1-a catalyst pipe to carry out a second regeneration reaction, and the second roasting is carried out in an air atmosphere to obtain a regenerated DPC1-a catalyst; wherein the temperature of the second roasting is 700 ℃, the pressure of the second roasting is 0.27MPa, the linear velocity of dense-phase bed flue gas of the second regenerator is 0.5m/s, and the linear velocity of dilute-phase bed flue gas is 0.55m/s; the obtained regenerated DPC1-a catalyst is circulated back to the downer reaction tube through the DPC1-a catalyst feeding tube;

(6) Introducing the reaction oil gas into a fractionating tower for fractionating to obtain dry gas, liquefied gas, gasoline, diesel oil and heavy fraction, and returning the obtained heavy fraction as recycle oil to a second reaction zone for recycling; wherein, the yield of the dry gas is 3.7%, the yield of the liquefied gas is 28.3%, the yield of the gasoline is 29.3%, the yield of the diesel is 23.3%, the yield of the heavy fraction is 8.0%, and the yield of the coke is 7.4%; wherein the ethylene content in the dry gas is 45.9wt%, the total content of propylene and butene in the liquefied gas is 64.4wt%, and the olefin content in the gasoline is 10.2wt%.

Example 2

(1) DPC1-a catalyst from the second regenerator enters a downer reaction tube through a DPC1-a catalyst feed tube, raw oil (intermediate-naphthenic coking wax oil) enters the downer reaction tube through a raw oil inlet, fluidization medium (water vapor) enters the downer reaction tube through a fluidization medium inlet, the DPC1-a catalyst contacts with the raw oil in the downer reaction tube under the carrying of the fluidization medium, flows from top to bottom in a first reaction zone, and a first reaction occurs to obtain a first mixture; wherein the feed temperature of the raw oil is 220 ℃, the feed temperature of the DPC1-a catalyst is 600 ℃, and the feed temperature of the fluidization medium is 275 ℃; the mass ratio of the DPC1-a catalyst to the raw oil is 8:1, the mass ratio of the raw oil to the fluidization medium is 100:1, a step of; the temperature of the first reaction is 480 ℃, the pressure of the first reaction is 0.23MPa, and the time of the first reaction is 2.0s;

(2) The pre-lifting medium (water vapor) enters a lifting pipe through a pre-lifting medium inlet, the DPC2-a catalyst from the first regenerator enters the lifting pipe through a DPC2-a catalyst feeding pipe, the first mixture is contacted with the DPC2-a catalyst and recycle oil under the carrying of the pre-lifting medium, and flows from bottom to top in a second reaction zone to generate a second reaction to obtain a second mixture; wherein the recycle oil is a heavy fraction, the feed temperature of the recycle oil is 250 ℃, and the feed temperature of the DPC2-a catalyst is 660 ℃; the feed temperature of the pre-lifting medium was 250 ℃, the mass ratio of recycle oil to feed oil (recycle ratio) was 0.25, and the mass ratio of feed oil to pre-lifting medium was 100:2.5, the DPC2-a catalyst was used in such an amount that the temperature of the second reaction was 510℃and the pressure of the second reaction was 0.23MPa, the time of the second reaction was 3.0s,

(3) The second mixture obtained by the riser reaction enters a settler to be contacted with stripping steam for separation, and a spent mixed catalyst and reaction oil gas are obtained; wherein, the mass ratio of the raw oil to the stripping steam is 100:4.5;

(4) The separated spent mixed catalyst enters a first regenerator through a spent mixed catalyst pipe to carry out a first regeneration reaction, and is subjected to first roasting under the air atmosphere to obtain a regenerated DPC2-a catalyst and a semi-regenerated DPC1-a catalyst; wherein the temperature of the first roasting is 680 ℃, the pressure of the first roasting is 0.27MPa, the dense-phase bed flue gas linear velocity of the first regenerator is 0.8m/s, and the dilute-phase bed flue gas linear velocity is 1.5m/s; the obtained regenerated DPC2-a catalyst is circulated back to the riser through a DPC2-a catalyst feed pipe;

(5) The obtained semi-regenerated DPC1-a catalyst enters a second regenerator through a semi-regenerated DPC1-a catalyst pipe to carry out a second regeneration reaction, and the second roasting is carried out in an air atmosphere to obtain a regenerated DPC1-a catalyst; wherein the temperature of the second roasting is 700 ℃, the pressure of the second roasting is 0.27MPa, the linear velocity of dense-phase bed flue gas of the second regenerator is 0.55m/s, and the linear velocity of dilute-phase bed flue gas is 0.6m/s; the obtained regenerated DPC1-a catalyst is circulated back to the downer reaction tube through the DPC1-a catalyst feeding tube;

(6) Introducing the reaction oil gas into a fractionating tower for fractionation to obtain dry gas, liquefied gas, gasoline, diesel oil and heavy fraction, and taking the obtained heavy fraction as recycle oil to return to a riser for recycling; wherein, the yield of the dry gas is 4.4%, the yield of the liquefied gas is 32.5%, the yield of the gasoline is 25.0%, the yield of the diesel is 22.0%, the yield of the heavy fraction is 9.1%, and the yield of the coke is 7.0%; wherein the ethylene content in the dry gas is 52.3wt%, the total content of propylene and butylene in the liquefied gas is 68.6wt%, and the content of olefin in the gasoline is 11.4wt%.

The preferred embodiments of the present invention have been described in detail above, but the present invention is not limited thereto. Within the scope of the technical idea of the invention, a number of simple variants of the technical solution of the invention are possible, including combinations of the individual technical features in any other suitable way, which simple variants and combinations should likewise be regarded as being disclosed by the invention, all falling within the scope of protection of the invention.

Claims

1. A method for catalytic cracking of a feedstock, said method comprising the steps of:

2. The method of claim 1, wherein in step (1), the feedstock oil is selected from crude oil and/or heavy oil;

preferably, the DPC1-a catalyst comprises 85-99 parts by weight of carrier I and 1-15 parts by weight of active metal oxide I; wherein the carrier I is at least one selected from alumina, silica, titania and zirconia; the active metal oxide I is selected from alkali metal oxide and/or alkaline earth metal oxide;

preferably, the DPC1-a catalyst is CO ₂ The desorption peak temperature is 185-195 ℃, preferably 187-192 ℃; CO ₂ The number of the alkali centers at the desorption peak positions is 16-22mmol/g, preferably 18-21mmol/g;

preferably, the DPC1-a catalyst has a bulk ratio of 0.5-0.7g/mL, preferably 0.55-0.65g/mL; particle size of 30-100 μm, preferably 50-70 μm;

preferably, the fluidizing medium is selected from at least one of steam, dry gas, natural gas and liquefied gas, preferably steam;

preferably, the feed temperature of the feedstock is 180-340 ℃, preferably 200-330 ℃; the feed temperature of the DPC1-a catalyst is 580-750 ℃, preferably 600-670 ℃; the feed temperature of the fluidization medium is 200-330 ℃, preferably 250-300 ℃;

Preferably, the mass ratio of the DPC1-a catalyst to the raw oil is 5-20:1, preferably 8-15:1, a step of; the mass ratio of the raw oil to the fluidization medium is 100:1-15, preferably 100:1-8;

preferably, the temperature of the first reaction is 370-535 ℃, preferably 410-525 ℃; the pressure of the first reaction is 0.1-1MPa, preferably 0.1-0.4MPa; the time of the first reaction is 0.1 to 4s, preferably 0.5 to 3s.

3. A method according to claim 1 or 2, wherein in step (2) the pre-lifting medium is selected from at least one of steam, dry gas, natural gas and liquefied gas, preferably steam;

preferably, the DPC2-a catalyst comprises 54-63 parts by weight of carrier II,1-5 parts by weight of active metal oxide II and 33-43 parts by weight of molecular sieve; wherein the carrier II is at least one of alumina, silica, titania and zirconia; the active metal oxide II is selected from alkali metal oxide and/or alkaline earth metal oxide; the molecular sieve comprises a medium pore molecular sieve and a large pore molecular sieve, wherein the medium pore molecular sieve is selected from ZSM-5 molecular sieve and/or ZRP-1 molecular sieve, and the large pore molecular sieve is selected from REHY type molecular sieve and/or REUSY type molecular sieve;

Preferably, the mass ratio of the medium pore molecular sieve to the large pore molecular sieve is 100:3-15, preferably 100:5-10;

preferably, the medium pore molecular sieve is selected from ZSM-5 molecular sieve and ZRP-1 molecular sieve, and the large pore molecular sieve is REUSY type molecular sieve;

preferably, the DPC2-a catalyst is CO ₂ The desorption peak temperature is 160-190 ℃, preferably 170-185 ℃; CO ₂ The number of the alkali centers at the desorption peak position is 5-18mmol/g, preferably 7-15mmol/g;

preferably, the DPC2-a catalyst has a bulk ratio of 0.65-0.95g/mL, preferably 0.75-0.85g/mL; particle size of 40-120 μm, preferably 70-90 μm;

preferably, the DPC1-a catalyst is CO ₂ The number of base centers at the desorption peak position is at least 10mmol/g more than that of the DPC2-a catalyst; the DPC2-a catalyst has a bulk ratio at least 0.15-0.25g/mL greater than the DPC1-a catalyst;

preferably, the cycle oil is a heavy fraction from the reaction oil gas;

preferably, the feed temperature of the cycle oil is 200-350 ℃, preferably 250-330 ℃; the feed temperature of the DPC2-a catalyst is 660-750 ℃, preferably 680-720 ℃; the feed temperature of the pre-lifting medium is 200-300 ℃, preferably 230-280 ℃;

Preferably, the mass ratio of recycle oil to feed oil (recycle ratio) is from 0.1 to 0.5, preferably from 0.2 to 0.4; the mass ratio of the raw oil to the pre-lifting medium is 100:1-15, preferably 100:1-8;

preferably, the reaction conditions of the second reaction further include: the DPC2-a catalyst is used in an amount such that the reaction temperature of the second reaction is 500-560 ℃, preferably 515-530 ℃; the pressure of the second reaction is 0.1-0.5MPa, preferably 0.2-0.3MPa; the time for the second reaction is 0.5 to 5s, preferably 1 to 3s.

4. A process according to any one of claims 1 to 3, wherein in step (3) the second mixture is separated by contact with stripping steam, wherein the mass ratio of feedstock oil to stripping steam is 100:1-20, preferably 100:1-10.

5. The method of any one of claims 1-4, wherein in step (4), the first regeneration reaction comprises a first calcination under an air atmosphere; preferably, the temperature of the first firing is 650-730 ℃, preferably 680-720 ℃; the pressure of the first roasting is 0.1-0.5MPa, preferably 0.15-0.35MPa; the dense bed flue gas line velocity of the first regenerator is 0.3-1m/s, preferably 0.5-0.9m/s; the linear velocity of the flue gas of the dilute phase bed in the first regenerator is 0.8-1.8m/s, preferably 1.1-1.5m/s;

Preferably, in step (5), the second regeneration reaction comprises performing a second calcination under an air and/or water vapor atmosphere; preferably, the temperature of the second firing is 650-720 ℃, preferably 660-700 ℃; the pressure of the second roasting is 0.1-0.5MPa, preferably 0.15-0.35MPa; the dense bed flue gas linear velocity of the second regenerator is 0.3-0.9m/s, preferably 0.5-0.8m/s; the flue gas linear velocity of the dilute phase bed of the second regenerator is 0.4-1.1m/s, preferably 0.5-0.9m/s.

6. The process according to any one of claims 1 to 5, wherein in step (6), the gasoline has a distillation range of 30 to 205 ℃, the diesel has a distillation range of 206 to 380 ℃ and the heavy fraction is a fraction oil having a temperature of 380 ℃ or higher.

7. The catalytic cracking device is characterized by comprising a descending bed reaction tube (1), a lifting tube (2), a settler (3), a first regenerator (4) and a second regenerator (5) which are connected end to end in sequence.

8. Catalytic cracking unit according to claim 7, wherein the second regenerator (5) is connected to the downer reaction tube (1) by a DPC1-a catalyst feed tube;

preferably, a raw oil inlet is arranged on the downer reaction tube (1); wherein the raw oil inlet is arranged below the joint of the DPC1-a catalyst feeding pipe and the downer reaction pipe (1);

Preferably, the downer reaction tube (1) is further provided with a fluidization medium inlet, wherein the fluidization medium inlet is arranged above the connection part of the DPC1-a catalyst feeding tube and the downer reaction tube (1).

9. Catalytic cracking unit according to claim 7 or 8, wherein the lower end of the downer reaction tube (1) is connected to the lower end of the riser (2), the riser (2) and the downer reaction tube (1) being "U" -shaped;

preferably, a DPC2-a catalyst inlet and a recycle oil inlet are arranged at the lower end of the riser (2); further preferably, the DPC2-a catalyst inlet and cycle oil inlet are arranged above the connection of the riser (2) and the downer reaction tube (1);

preferably, the lower end of the lifting pipe is also provided with a pre-lifting medium inlet; wherein the lifting medium inlet is arranged below the joint of the downer reaction tube (1) and the lifting tube (2);

preferably, the upper end of the riser (2) is connected to a settler (3).

10. Catalytic cracking unit according to any one of claims 7-9, wherein the settler (3) is connected to the first regenerator (4) via a spent mixed catalyst feed pipe, the first regenerator (4) being connected to the riser via a DPC2-a catalyst feed pipe;

Preferably, the first regenerator (4) is also connected to the second regenerator (5) by a semi-regenerated DPC1-a catalyst feed line;

preferably, a material level adjusting pipe is arranged between the second regenerator (5) and the first regenerator (4), and is used for adjusting the material level in the second regenerator (5);

preferably, the device further comprises a fractionating tower connected with the settler (3), wherein the fractionating tower is used for separating the reaction oil gas to obtain dry gas, liquefied gas, gasoline, diesel oil and heavy fraction.