CN111100710B

CN111100710B - Catalytic cracking method and system

Info

Publication number: CN111100710B
Application number: CN201811251019.8A
Authority: CN
Inventors: 李泽坤; 龚剑洪; 唐津莲; 毛安国
Original assignee: Sinopec Research Institute of Petroleum Processing; China Petroleum and Chemical Corp
Current assignee: Sinopec Research Institute of Petroleum Processing; China Petroleum and Chemical Corp
Priority date: 2018-10-25
Filing date: 2018-10-25
Publication date: 2022-03-11
Anticipated expiration: 2038-10-25
Also published as: CN111100710A

Abstract

The present invention relates to a method and system for catalytic cracking, the method comprising: a. carrying out catalytic cracking reaction on a heavy raw material in the presence of a first catalytic cracking catalyst to obtain a first reaction product and a first catalyst to be generated; b. sending the obtained first reaction product into a product separation device for separation to obtain at least catalytic cracking gasoline and catalytic cracking light cycle oil; c. cutting the catalytic cracking light cycle oil to obtain light fraction and heavy fraction; wherein the cut points of the light fraction and the heavy fraction are in the range of 240 ℃ and 260 ℃; d. introducing the obtained heavy fraction into a hydrogenation reactor to contact with a hydrotreating catalyst and carrying out hydrotreating to obtain hydrogenated heavy fraction; e. and carrying out catalytic cracking reaction on the obtained hydrogenated heavy fraction and light fraction in the presence of a second catalytic cracking catalyst to obtain a second reaction product and a second spent catalyst. The method and the system can produce gasoline with high octane number.

Description

Catalytic cracking method and system

Technical Field

The present invention relates to a catalytic cracking process and system.

Background

With the development of crude oil upgrading and the rapid increase in the market demand for light oils, catalytic cracking technology for upgrading heavy oils is rapidly developed in china. However, one must face the fact that catalytically cracked diesel (or light cycle oil) has been relatively poor in quality, high in density, high in aromatics content, low in cetane number, and even by diesel hydro-upgrading technology, it has been difficult to meet increasingly stringent diesel specifications. How to treat the catalytic cracking light cycle oil is an increasingly serious problem. Meanwhile, another problem exists in that domestic finished gasoline is in short supply for a long time, and catalytic cracking gasoline accounts for 80 percent of the finished gasoline. Therefore, how to achieve maximum production of high octane gasoline from heavy feedstock without producing light cycle oil by catalytic cracking process may be a new way to solve the above problems.

US patent US4585545 discloses a catalytic conversion method for producing gasoline rich in monocyclic aromatic hydrocarbons by hydrotreating a whole fraction of catalytically cracked light cycle oil to obtain hydrogenated diesel oil and then catalytically cracking the hydrogenated diesel oil.

Chinese patent application CN14232327A discloses a method for upgrading catalytic cracking cycle oil, which is to deeply hydrogenate light cycle oil produced by a first catalytic cracking unit using heavy oil as a raw material, and then to subject the obtained hydrogenated cycle oil to a second catalytic cracking unit. On the basis of this process, chinese patent application publication CN423689A emphasizes that the catalyst in the second catalytic cracking unit requires 50-95% shape selective zeolite and 5-50% large pore zeolite with a pore size greater than or equal to about 0.7nm to selectively increase the light olefin yield.

Chinese patent application CN1466619A discloses a conversion method of catalytic cracking light cycle oil, which is to divide a catalytic cracking riser reactor into an upstream reaction zone and a downstream reaction zone, wherein heavy oil is injected into the downstream reaction zone, and hydrogenated cycle oil obtained by hydrotreating the catalytic cracking product light cycle oil is injected into the upstream reaction zone. On the basis of the method, the feed of the upstream zone in the method disclosed in the Chinese patent application publication CN1425054A is added with naphtha in addition to the hydrogenated cycle oil. However, in the method, not only is the hydrogen consumption high, but also the hydrogenation cycle oil reacts in an upstream area to seriously affect the conversion of heavy oil in the downstream area.

There remains a need in the art for a process that reduces the yield of light cycle oil, increases the yield of catalytically cracked gasoline, and simultaneously increases its octane number.

Disclosure of Invention

The invention aims to provide a catalytic cracking method and a catalytic cracking system, which can produce high-octane gasoline in a high yield.

In order to achieve the above object, the present invention provides a catalytic cracking process comprising:

a. carrying out catalytic cracking reaction on a heavy raw material in the presence of a first catalytic cracking catalyst to obtain a first reaction product and a first catalyst to be generated;

b. sending the obtained first reaction product into a product separation device for separation to obtain at least catalytic cracking gasoline and catalytic cracking light cycle oil;

c. cutting the catalytic cracking light cycle oil to obtain light fraction and heavy fraction; wherein the cut points of the light fraction and the heavy fraction are in the range of 240 ℃ and 260 ℃;

d. introducing the obtained heavy fraction into a hydrogenation reactor to contact with a hydrotreating catalyst and carrying out hydrotreating to obtain hydrogenated heavy fraction;

e. carrying out catalytic cracking reaction on the obtained hydrogenated heavy fraction and light fraction in the presence of a second catalytic cracking catalyst to obtain a second reaction product and a second spent catalyst;

the catalytic cracking reaction in the step a and the catalytic cracking reaction in the step e are carried out in a first riser reactor, the heavy raw material and the hydrogenated heavy fraction are respectively sprayed into the first riser reactor from a first nozzle and a second nozzle which are arranged from bottom to top at intervals and react from bottom to top, the light fraction is sprayed into the first riser reactor from a third nozzle, the introducing position of the first catalytic cracking catalyst is arranged at the bottom of the first riser reactor and along the height direction of the first riser reactor, and the introducing position of the second catalytic cracking catalyst is arranged between the first nozzle and the second nozzle; or

And c, respectively carrying out the catalytic cracking reaction in the step a and the catalytic cracking reaction in the step e in a second riser reactor and a third riser reactor, wherein the top opening of the third riser reactor is communicated with the middle upper part of the second riser reactor through a horizontal pipe, and the second reaction product and the second spent catalyst are fed into the middle upper part of the second riser reactor from the top opening of the third riser reactor and are discharged from the top opening of the second riser reactor together with the first reaction product and the first spent catalyst.

Optionally, along the height direction of the first riser reactor, the third nozzle is positioned between the first nozzle and the second nozzle, and the introduction position of the second catalytic cracking catalyst is positioned between the first nozzle and the third nozzle;

and the hydrogenated heavy fraction is sprayed into the third riser reactor from a fourth nozzle, the light fraction is sprayed into the third riser reactor from a fifth nozzle, and the fourth nozzle is positioned above the fifth nozzle along the height direction of the third riser reactor.

Optionally, in the first riser reactor between the first nozzle and the second nozzle, the residence time of the reaction oil gas is 0.05 to 3 seconds; in the first riser reactor between the first nozzle and the third nozzle, the residence time of the reaction oil gas is 0.01-2 seconds;

and in the second riser reactor between the fourth nozzle and the fifth nozzle, the residence time of the reaction oil gas is 0.01-2 seconds.

Optionally, in the first riser reactor, the circulating weight ratio of the first catalytic cracking catalyst to the second catalytic cracking catalyst is 1: (0.02-1), the circulating weight ratio of the light fraction is 0.01-1, and the circulating weight ratio of the heavy fraction is 0.01-1;

in the second riser reactor and the third riser reactor, the circulating weight ratio of the light fraction is 0.01-1, and the circulating weight ratio of the heavy fraction is 0.01-1.

Optionally, the reaction conditions of the first riser reactor include: the reaction temperature is 520 ℃ and 650 ℃, the absolute pressure is 0.15-0.4 MPa, the weight ratio of the total weight of the first catalytic cracking catalyst and the second catalytic cracking catalyst to the heavy raw material is 1-50, the oil gas retention time of the heavy raw material is 1-10 seconds, and the weight ratio of the water vapor to the heavy raw material is 0.01-0.5; the proportion of the total weight of the first catalytic cracking catalyst and the second catalytic cracking catalyst to the total weight of the hydrogenated heavy fraction and the light fraction is 5-100, the proportion of the weight of the steam to the total weight of the hydrogenated heavy fraction and the light fraction is 0.01-0.3, and the micro-reverse activities of the first catalytic cracking catalyst and the second catalytic cracking catalyst are not less than 60;

the reaction conditions of the second riser reactor include: the reaction temperature is 450-550 ℃, the weight ratio of the first catalytic cracking catalyst to the heavy raw material is 4-8, the oil gas retention time is 2-10 seconds, the absolute pressure is 0.15-0.4 MPa, the weight ratio of the water vapor to the heavy raw material is 0.02-0.08, and the micro-reaction activity of the first catalytic cracking catalyst is not lower than 60;

the reaction conditions of the third riser reactor include: the reaction temperature is 520 ℃ and 650 ℃, the absolute pressure is 0.15-0.4 MPa, the ratio of the weight of the second catalytic cracking catalyst to the total weight of the hydrogenated heavy fraction and the light fraction is 5-100, the retention time of oil gas of the hydrogenated heavy fraction is 1-10 seconds, the ratio of the weight of the water vapor to the total weight of the hydrogenated heavy fraction and the light fraction is 0.01-0.3, and the micro-reaction activity of the second catalytic cracking catalyst is not lower than 60.

Optionally, the first riser reactor, the second riser reactor and the third riser reactor are respectively an equal-diameter riser reactor or a reducing riser reactor.

Optionally, the second riser reactor is coaxially provided with a first reaction section and a second reaction section from bottom to top, the inner diameter of the second reaction section is larger than that of the first reaction section, and the top opening of the third riser reactor is communicated with the second reaction section through a horizontal pipe.

Optionally, the heavy feedstock is selected from one or more of straight run wax oil, coker wax oil, deasphalted oil, hydrofined oil, hydrocracked tail oil, vacuum residuum, and atmospheric residuum.

Optionally, the first and second catalytic cracking catalysts each comprise 10-50 wt% of a zeolite selected from one or more of rare earth-containing or non-containing Y-type zeolite, HY-type zeolite, USY-type zeolite and Beta zeolite, 5-90 wt% of an inorganic oxide selected from silica and/or aluminum trioxide, and 0-70 wt% of a clay selected from kaolin and/or halloysite.

Optionally, the hydrotreating catalyst includes an active metal component and a carrier, the active metal component is a group VIB metal and/or a group VIII non-noble metal, and the carrier is at least one selected from alumina, silica, and amorphous silica-alumina.

Optionally, the active metal component is nickel-tungsten, nickel-tungsten-cobalt, nickel-molybdenum or cobalt-molybdenum.

Optionally, the hydrotreating conditions include: hydrogen partial pressure of 5.0-22.0 MPa, reaction temperature of 330-^-1Hydrogen oil volume ratio of 100-2000Nm³/m³。

Optionally, the content of bicyclic aromatic hydrocarbons in the hydrogenated heavy fraction is not more than 20 wt%, and the content of hydrogen is not less than 10 wt%.

The invention also provides a catalytic cracking system, which comprises a reaction unit, a regenerator, a hydrogenation reactor, a settler and a product separation device;

the product separation device is provided with a raw material inlet, a catalytic cracking gasoline outlet, a light fraction outlet and a heavy fraction outlet, the hydrogenation reactor is provided with a raw material inlet and a hydrogenation heavy fraction outlet, and the heavy fraction outlet of the product separation device is communicated with the raw material inlet of the hydrogenation reactor;

wherein the reaction unit comprises a first riser reactor, the first riser reactor is provided with a first catalytic cracking catalyst inlet positioned at the bottom, a second catalytic cracking catalyst inlet positioned at the middle part, a first nozzle used for feeding heavy raw materials, a second nozzle used for feeding hydrogenated heavy fractions, a third nozzle used for feeding light fractions and a top opening used for feeding oil, the first nozzle is positioned above the second nozzle, the second nozzle is communicated with a hydrogenated heavy fraction outlet of the hydrogenation reactor, the third nozzle is communicated with a light fraction outlet of the product separation device, the top opening of the first riser reactor is communicated with the settler, a catalyst outlet of the settler is communicated with a catalyst inlet of the regenerator, and an oil-gas outlet of the settler is communicated with a raw material inlet of the product separation device, the catalyst outlet of the regenerator is communicated with the first catalytic cracking catalyst inlet and the second catalytic cracking catalyst inlet of the first riser reactor; or

The reaction unit comprises a second riser reactor and a third riser reactor, the second riser reactor is provided with a heavy raw material inlet at the lower part, a catalytic cracking catalyst inlet at the bottom and a top opening for sending out oil, the second riser reactor is provided with a fourth nozzle for sending in hydrogenated heavy fraction, a fifth nozzle for sending in light fraction, a catalytic cracking catalyst inlet at the bottom and a product outlet at the top, the fourth nozzle is communicated with the hydrogenated heavy fraction outlet of the hydrogenation reactor, the fifth nozzle is communicated with the light fraction outlet of the product separation device, the top opening of the third riser reactor is communicated with the middle upper part of the second riser reactor through a horizontal pipe, the top opening of the second riser reactor is communicated with the settler, and the catalyst outlet of the settler is communicated with the catalyst inlet of the regenerator, the oil gas outlet of the settler is communicated with the raw material inlet of the product separation device, and the catalyst outlet of the regenerator is communicated with the catalytic cracking catalyst inlets of the second riser reactor and the third riser reactor.

Optionally, along the height direction of the first riser reactor, the third nozzle is positioned between the first nozzle and the second nozzle, and the second catalytic cracking catalyst inlet is positioned between the first nozzle and the third nozzle;

the fourth nozzle is positioned above the fifth nozzle in the height direction of the third riser reactor.

The method and system of the present invention can achieve one or more of the following advantages over prior art methods:

1. the light cycle oil is not generated thoroughly.

2. The heavy raw material, the hydrogenated heavy fraction and the light fraction are respectively processed by feeding materials at different height positions of the same riser reactor and supplementing a catalytic cracking catalyst, so that the operation conditions of the three raw materials are respectively optimized, the maximum conversion of the three raw materials is realized, and the catalytic cracking gasoline with high octane number is produced more. Furthermore, the hydrogenated heavy fraction and the hydrogenated light fraction are fed at the downstream of the heavy raw material, so that the reaction time can be effectively shortened, and the yield of the high-octane gasoline can be further improved.

3. The heavy raw material, the hydrogenated heavy fraction and the light fraction are respectively processed by adopting the second riser reactor and the third riser reactor, different operation parameters can be adopted in the second riser reactor and the third riser reactor, the harsh conditions required by the catalytic cracking of the hydrogenated heavy fraction and the light fraction are optimized and met to the maximum extent, and meanwhile, the top outlet of the third riser reactor is communicated with the middle upper part of the second riser reactor through a horizontal pipe, so that the reaction time of the hydrogenated heavy fraction and the light fraction can be shortened, and the catalytic cracking gasoline with high octane number can be produced more; meanwhile, the device structure is simplified, and the equipment cost is reduced.

Additional features and advantages of the invention will be set forth in the detailed description which follows.

Drawings

The accompanying drawings, which are included to provide a further understanding of the invention and are incorporated in and constitute a part of this specification, illustrate embodiments of the invention and together with the description serve to explain the principles of the invention and not to limit the invention. In the drawings:

FIG. 1 includes a schematic flow diagram of one embodiment of the method of the present invention and also includes a schematic structural diagram of one embodiment of the system of the present invention.

FIG. 2 comprises a schematic flow diagram of another embodiment of the method of the present invention, and also comprises a schematic structural diagram of another embodiment of the system of the present invention.

Description of the reference numerals

201 heavy fraction line 202 hydrogen line 203 hydrogenation reactor

204 hydrogenated heavy ends line 205 heavy feed line 206 second nozzle

207 third nozzle 208 first nozzle 209 second regeneration chute

210 second regenerative spool valve 211 first regenerative ramp 212 first regenerative spool valve

213 settler 214 regenerator 215 product line

216 oil gas pipeline 217 main fractionating tower 218 light cycle oil fractionating tower

219 light ends line 220 recycle line 221 first riser reactor

401 second riser reactor 402 light cycle oil fractionation column 403 main fractionation column

404 third riser reactor 405 regenerator 406 settler

407 sixth nozzle 408 fourth nozzle 409 second regeneration chute

410 first regeneration inclined tube 411 light cycle oil pipeline 412 hydrogenation heavy fraction pipeline

413 heavy fraction line 414 oil and gas line 415 slurry line

416 hydrogen line 417 product line 418 light ends line

419 fifth nozzle 420 hydrogenation reactor

I a first reaction section II a second reaction section

Detailed Description

The following describes in detail specific embodiments of the present invention. It should be understood that the detailed description and specific examples, while indicating the present invention, are given by way of illustration and explanation only, not limitation.

In a first aspect the present invention provides a process for catalytic cracking, the process comprising:

the catalytic cracking reaction in the step a and the catalytic cracking reaction in the step e are carried out in a first riser reactor, the heavy raw material and the hydrogenated heavy fraction are respectively sprayed into the first riser reactor from a first nozzle and a second nozzle which are arranged from bottom to top at intervals and react from bottom to top, the light fraction is sprayed into the first riser reactor from a third nozzle, the introducing position of the first catalytic cracking catalyst is arranged at the bottom of the first riser reactor and along the height direction of the first riser reactor, and the introducing position of the second catalytic cracking catalyst is arranged between the first nozzle and the second nozzle.

According to the first aspect of the invention, the cut point of the light fraction and the heavy fraction in the range of 240 ℃ and 260 ℃ means any point value between the cut point of 240 ℃ and 260 ℃, such as 240 ℃, 245 ℃, 250 ℃ or 260 ℃ and the like.

According to the first aspect of the present invention, the method may specifically include:

spraying a heavy raw material into a first riser reactor from a first nozzle, and carrying out catalytic cracking reaction in the presence of a first catalytic cracking catalyst and a second catalytic cracking catalyst to obtain a first reaction product and a first catalyst to be generated;

sending the obtained first reaction product into a product separation device for separation to obtain at least catalytic cracking gasoline and catalytic cracking light cycle oil;

cutting the catalytic cracking light cycle oil to obtain light fraction and heavy fraction; wherein the cut points of the light fraction and the heavy fraction are in the range of 240 ℃ and 260 ℃;

introducing the obtained heavy fraction into a hydrogenation reactor to contact with a hydrotreating catalyst and carrying out hydrotreating to obtain hydrogenated heavy fraction;

spraying the obtained hydrogenated heavy fraction and light fraction into the first riser reactor from a second nozzle and a third nozzle respectively, and carrying out catalytic cracking reaction from bottom to top to obtain a second reaction product and a second spent catalyst; the first nozzle and the second nozzle are arranged at intervals from bottom to top in the first riser reactor, the introduction position of the first catalytic cracking catalyst is arranged at the bottom of the first riser reactor, and the introduction position of the second catalytic cracking catalyst is arranged between the first nozzle and the second nozzle along the height direction of the first riser reactor.

According to the first aspect of the present invention, the third nozzle may be positioned between the first nozzle and the second nozzle, and the introduction position of the second catalytic cracking catalyst may be positioned between the first nozzle and the third nozzle, along the height direction of the first riser reactor.

According to the first aspect of the present invention, in the first riser reactor between the first nozzle and the second nozzle, the residence time of the reaction oil gas may be 0.05 to 3 seconds, preferably 0.1 to 1 second; in the first riser reactor between the first nozzle and the third nozzle, the residence time of the reaction oil gas can be 0.01-2 seconds.

According to the first aspect of the present invention, a second catalytic cracking catalyst is supplemented in the middle of the first riser reactor to enhance the conversion of the hydrogenated heavy fraction and light fraction, the first catalytic cracking catalyst enters the riser reactor from a conventional catalyst inlet of the riser reactor, and the second catalytic cracking catalyst is fed from a supplement inlet in the middle of the riser reactor. In the first riser reactor, the circulating weight ratio of the first catalytic cracking catalyst to the second catalytic cracking catalyst per unit time may be 1: (0.02-1), preferably 1: (0.03-0.5), the recycle weight ratio of the light fraction can be 0.01-1, and the recycle weight ratio of the heavy fraction can be 0.01-1;

according to the first aspect of the present invention, the reaction conditions of the riser reactor are well known to those skilled in the art, and in particular for the present invention, the reaction conditions of the first riser reactor may include: the reaction temperature is 520-650 ℃, preferably 550-590 ℃, the absolute pressure is 0.15-0.4 MPa, the weight ratio of the total weight of the first catalytic cracking catalyst and the second catalytic cracking catalyst to the heavy raw material is 1-50, preferably 3-30, the oil gas residence time of the heavy raw material is 1-10 seconds, preferably 2-8 seconds, and the weight ratio of the water vapor to the heavy raw material is 0.01-0.5, preferably 0.02-0.2; the ratio of the total weight of the first catalytic cracking catalyst and the second catalytic cracking catalyst to the total weight of the hydrogenated heavy fraction and the light fraction is 5-100, preferably 8-50, the ratio of the weight of the steam to the total weight of the hydrogenated heavy fraction and the light fraction is 0.01-0.3, preferably 0.02-0.1, and the micro-reverse activities of the first catalytic cracking catalyst and the second catalytic cracking catalyst are not less than 60, preferably not less than 62.

According to the first aspect of the present invention, the first riser reactor may be a constant diameter riser reactor or a variable diameter riser reactor.

In a second aspect the present invention provides a process for catalytic cracking, the process comprising:

According to the second aspect of the invention, the cut point of the light fraction and the heavy fraction in the range of 240-260 ℃ means any point value between the cut point of 240-260 ℃, such as 240 ℃, 245 ℃, 250 ℃ or 260 ℃ and the like.

According to the second aspect of the present invention, the method may specifically include:

introducing a heavy raw material into a second riser reactor to contact with a first catalytic cracking catalyst and carrying out catalytic cracking reaction to obtain a first reaction product and a first catalyst to be generated;

introducing the obtained hydrogenated heavy fraction and light fraction into a third riser reactor to contact with a second catalytic cracking catalyst and perform catalytic cracking reaction to obtain a second reaction product and a second spent catalyst; the top opening of the third riser reactor is communicated with the middle upper part of the second riser reactor through a horizontal pipe, and the second reaction product and the second spent catalyst are sent into the middle upper part of the second riser reactor from the top opening of the third riser reactor and are sent out from the top opening of the second riser reactor together with the first reaction product and the first spent catalyst.

According to a second aspect of the present invention, the hydrogenated heavy fraction is injected into the third riser reactor from a fourth nozzle, and the light fraction is injected into the third riser reactor from a fifth nozzle, and the fourth nozzle may be located above the fifth nozzle in the height direction of the third riser reactor, so as to facilitate the conversion of the light fraction. In the second riser reactor and the third riser reactor, the circulating weight ratio of the light fraction may be 0.01 to 1, and the circulating weight ratio of the heavy fraction may be 0.01 to 1. In the second riser reactor between the fourth nozzle and the fifth nozzle, the residence time of the reaction oil gas can be 0.01-2 seconds.

According to the second aspect of the present invention, the reaction conditions of the riser reactor are well known to those skilled in the art, and in particular for the present invention, the reaction conditions of the second riser reactor may include: the reaction temperature is 450-550 ℃, the weight ratio of the first catalytic cracking catalyst to the heavy raw material is 4-8, the oil gas retention time is 2-10 seconds, the absolute pressure is 0.15-0.4 MPa, the weight ratio of the water vapor to the heavy raw material is 0.02-0.08, and the micro-reaction activity of the first catalytic cracking catalyst is not lower than 60, preferably not lower than 62; the reaction conditions of the third riser reactor may include: the reaction temperature is 520-650 ℃, preferably 550-590 ℃, the absolute pressure is 0.15-0.4 MPa, the weight ratio of the second catalytic cracking catalyst to the total weight of the hydrogenated heavy fraction and the light fraction is 5-100, preferably 8-50, the oil gas retention time of the hydrogenated heavy fraction is 1-10 seconds, preferably 2-8 seconds, the weight ratio of the water vapor to the total weight of the hydrogenated heavy fraction and the light fraction is 0.01-0.3, preferably 0.02-0.2, and the micro-reaction activity of the second catalytic cracking catalyst is not less than 60, preferably not less than 62.

According to the second aspect of the present invention, the second riser reactor and the third riser reactor may each be a constant diameter riser reactor or a variable diameter riser reactor. The second riser reactor can be coaxially provided with a first reaction section and a second reaction section from bottom to top, the inner diameter of the second reaction section can be larger than that of the first reaction section, and the top opening of the third riser reactor can be communicated with the second reaction section through a horizontal pipe. If the height of the second riser reactor is h, the upper middle portion of the second riser reactor generally refers to a height of greater than 1/3h, preferably greater than 1/3h and less than 4/5.

Heavy oil feedstocks are well known to those skilled in the art in accordance with the first and second aspects of the present invention, and for example, the heavy feedstock may be at least one selected from straight run waxy oil, coker waxy oil, deasphalted oil, hydrofinished oil, hydrocracked tail oil, vacuum residue, and atmospheric residue, and other heavy feedstocks may be processed by those skilled in the art. The externally produced catalytic cracking light cycle oil can also be hydrogenated and then sent into a riser reactor as the hydrogenated light cycle oil for reaction, so that the source of raw materials is improved, and the high-octane gasoline is produced in more.

According to the first and second aspects of the present invention, the catalytic cracking catalyst is well known to those skilled in the art, for example, the first and second catalytic cracking catalysts may each comprise 10 to 50 wt% of a zeolite, which may be selected from one or more of rare earth-containing Y-type zeolite, HY-type zeolite, USY-type zeolite and Beta zeolite, 5 to 90 wt% of an inorganic oxide, which may be selected from silica and/or aluminum trioxide, and 0 to 70 wt% of a clay, which may be selected from kaolin and/or halloysite.

Hydrotreating according to the first and second aspects of the invention is well known to those skilled in the art, for example, the hydrotreating catalyst may comprise an active metal component and a support, preferably comprising 30 wt.% of the active metal component and 70 wt.% of the support. The active metal component can be a VIB group metal and/or VIII group non-noble metal, preferably nickel-tungsten, nickel-tungsten-cobalt, nickel-molybdenum or cobalt-molybdenum, and the carrier can be at least one selected from alumina, silica and amorphous silicon-aluminum. The hydrotreating conditions may include: partial pressure of hydrogen5.0-22.0 MPa, preferably 8.0-15.0 MPa, the reaction temperature is 330-^-1Preferably 0.1 to 3.0 hours^-1Hydrogen oil volume ratio of 100-2000Nm³/m³Preferably 350-2000Nm³/m³. The hydrogenated heavy fraction may have a bicyclic aromatic content of not more than 20 wt%, more preferably not more than 8 wt%, a hydrogen content of not less than 10 wt%, preferably not less than 11 wt%, more preferably not less than 14 wt%, and a first boiling point of the hydrogenated distillate may be greater than 165 ℃, preferably greater than 170 ℃, or the first boiling point may be greater than 250 ℃, preferably greater than 260 ℃.

The third aspect of the present invention also provides a system for catalytic cracking, comprising a reaction unit, a regenerator, a hydrogenation reactor, a settler and a product separation unit;

wherein the reaction unit comprises a first riser reactor, the first riser reactor is provided with a first catalytic cracking catalyst inlet positioned at the bottom, a second catalytic cracking catalyst inlet positioned at the middle part, a first nozzle used for feeding heavy raw materials, a second nozzle used for feeding hydrogenated heavy fractions, a third nozzle used for feeding light fractions and a top opening used for feeding oil, the first nozzle is positioned above the second nozzle, the second nozzle is communicated with a hydrogenated heavy fraction outlet of the hydrogenation reactor, the third nozzle is communicated with a light fraction outlet of the product separation device, the top opening of the first riser reactor is communicated with the settler, a catalyst outlet of the settler is communicated with a catalyst inlet of the regenerator, and an oil-gas outlet of the settler is communicated with a raw material inlet of the product separation device, and the catalyst outlet of the regenerator is communicated with the first catalytic cracking catalyst inlet and the second catalytic cracking catalyst inlet of the first riser reactor.

According to a third aspect of the present invention, the third nozzle may be positioned between the first nozzle and the second nozzle, and the second catalytic cracking catalyst inlet may be positioned between the first nozzle and the third nozzle, along a height direction of the first riser reactor.

According to the third aspect of the present invention, the first riser reactor may be a constant diameter riser reactor or a variable diameter riser reactor.

The fourth aspect of the present invention also provides a catalytic cracking system, comprising a reaction unit, a regenerator, a hydrogenation reactor, a settler and a product separation device;

wherein the reaction unit comprises a first riser reactor, the first riser reactor is provided with a first catalytic cracking catalyst inlet positioned at the bottom, a second catalytic cracking catalyst inlet positioned at the middle part, a first nozzle used for feeding heavy raw materials, a second nozzle used for feeding hydrogenated heavy fractions, a third nozzle used for feeding light fractions and a top opening used for feeding oil, the first nozzle is positioned above the second nozzle, the second nozzle is communicated with a hydrogenated heavy fraction outlet of the hydrogenation reactor, the third nozzle is communicated with a light fraction outlet of the product separation device, the top opening of the first riser reactor is communicated with the settler, a catalyst outlet of the settler is communicated with a catalyst inlet of the regenerator, and an oil-gas outlet of the settler is communicated with the product separation device, the catalyst outlet of the regenerator is communicated with the first catalytic cracking catalyst inlet and the second catalytic cracking catalyst inlet of the first riser reactor;

According to a fourth aspect of the present invention, the fourth nozzle may be positioned above the fifth nozzle in a height direction of the third riser reactor.

According to a fourth aspect of the present invention, the second riser reactor and the third riser reactor may each be a constant diameter riser reactor or a variable diameter riser reactor.

According to a fourth aspect of the present invention, the second riser reactor may be coaxially provided with a first reaction section and a second reaction section from bottom to top, the inner diameter of the second reaction section may be larger than that of the first reaction section, and the top opening of the third riser reactor may be communicated with the second reaction section through a horizontal pipe.

According to the third and fourth aspects of the invention, the product separation device is well known to those skilled in the art, and may be, for example, a fractionating tower, and the invention preferably comprises a main fractionating tower and a light cycle oil fractionating tower, wherein the main fractionating tower separates light cycle oil and catalytic cracking gasoline, and the light cycle oil is fed into the light cycle oil fractionating tower for further separation to obtain light fraction and heavy fraction.

The invention will be further illustrated by means of specific embodiments in the following description with reference to the drawings, but the invention is not limited thereto in any way.

According to the first and third aspects of the present invention, as shown in fig. 1, a heavy feedstock is introduced into the lower portion of the first riser reactor 221 through the first nozzle 208 via the heavy feedstock line 205, a heavy fraction cut by catalytically cracking light cycle oil is introduced into the hydrogenation reactor 203 via the heavy fraction line 201, and hydrogen is introduced into the hydrogenation reactor 203 via the hydrogen line 202. The hydrogenated heavy fraction obtained by hydrogenation enters the first riser reactor 221 through a hydrogenated heavy fraction pipeline 204 through a second nozzle 206, and the light fraction cut by the catalytic cracking light cycle oil enters the first riser reactor 221 through a light fraction pipeline 219 through a third nozzle 207. A part of the regenerated catalyst from the regenerator 214 passes through a first regeneration inclined tube 211, is controlled by a first regeneration slide valve 212, passes through the bottom of the first riser reactor 221 as a first catalytic cracking catalyst, ascends under the action of a pre-lifting medium, enters a heavy raw material reaction zone, and contacts, reacts and ascends with the heavy raw material injected through the first nozzle 208 in the reaction zone; the mixture of reacted oil and gas and catalyst then enters the cycle oil reaction zone where it contacts, reacts and travels upward with the hydrogenated heavy fraction fed through second nozzle 206 and the light fraction fed through third nozzle 207. Another portion of the regenerated catalyst from the regenerator 214 is injected into the first riser reactor 221 as a second catalytic cracking catalyst via a second regeneration ramp 209 and a second regeneration slide valve 210 between the first nozzle 208 and the second nozzle 206 to enhance the cracking reaction of the hydrogenated heavy fraction and the light fraction in the hydrocyclingoil reaction zone. The second nozzle 206, the third nozzle 207 and the first nozzle 208 are disposed in a relative positional relationship of upper-middle-lower along the height direction of the first riser reactor 221. The catalytic cracking reaction product and spent catalyst discharged from the top of the first riser reactor 221 enter a settler 213 for separation of the reaction product and the catalyst, and the separated spent catalyst enters a regenerator 214 for regeneration and circulation. The separated reaction product enters a main fractionating tower 217 through a product pipeline 215, and oil slurry discharged from the main fractionating tower 217 is taken as a product discharge device through an oil slurry pipeline; the oil gas from the main fractionating tower 217 enters a subsequent absorption stabilizing system (not shown in the figure) through an oil gas pipeline 216 to obtain dry gas, liquefied gas and high-octane gasoline; and the catalytic cracking light cycle oil from the main fractionating tower 217 is cut in the light cycle oil fractionating tower 218 to obtain light fraction and heavy fraction, the heavy fraction enters the hydrogenation reactor 203 for circulation through the recycle pipeline 220 and the heavy fraction pipeline 201, and the light fraction enters the first riser reactor 221 through the light fraction pipeline 219 and the third nozzle 207.

According to the second and fourth aspects of the present invention, as shown in fig. 2, the heavy feedstock enters the first reaction section I of the second riser reactor 401 through the sixth nozzle 407, the heavy fraction cut by the catalytically cracked light cycle oil enters the hydrogenation reactor 420 through the heavy fraction line 413, and hydrogen is introduced into the hydrogenation reactor 420 through the hydrogen line 416. The hydrogenated heavy fraction enters the third riser reactor 404 through the fourth nozzle 408 via the hydrogenated heavy fraction line 412, and the light fraction enters the third riser reactor 404 through the fifth nozzle 419 via the light fraction line 418. A portion of the regenerated catalyst from the regenerator 405 is injected into the bottom of the second riser reactor 401 through the first regeneration inclined tube 410 as the first catalytic cracking catalyst, ascends under the action of the pre-lift medium, contacts, reacts and ascends with the heavy feedstock injected through the sixth nozzle 407. Another part of the regenerated catalyst from the regenerator 305 is injected into the bottom of the third riser reactor 404 through a second regeneration inclined pipe 409 to serve as a second catalytic cracking catalyst, moves upwards under the action of the pre-lifting medium, contacts, reacts and moves upwards with the hydrogenated heavy fraction fed through a fourth nozzle 408 and the light fraction fed through a fifth nozzle 419, and the obtained reaction material is fed into a second reaction section II at the upper middle part of the second riser reactor 401 from the top outlet of the third riser reactor 404 through a horizontal pipe and is merged with the material therein. The fourth nozzle 408 and the fifth nozzle 419 are disposed in a relative positional relationship of up-down along the height direction of the third riser reactor 404. The catalytic cracking reaction product and spent catalyst discharged from the top of the second riser reactor 401 enter a settler 406 for separation of the reaction product and the catalyst, and the separated spent catalyst enters a regenerator 405 for regeneration and circulation. The separated reaction product enters the fractionating tower 403 through a product pipeline 417, and the oil slurry from the fractionating tower 403 is discharged out of the device as a product through an oil slurry pipeline 415; the oil gas from the fractionating tower 403 enters a subsequent absorption stabilizing system (not shown in the figure) through an oil gas pipeline 414 to obtain dry gas, liquefied gas and high-octane gasoline; and the catalytic cracking light cycle oil from the main fractionating tower 403 is sent to the light cycle oil fractionating tower 402 through a light cycle oil pipeline 411, and is cut to obtain light fraction and heavy fraction, the catalytic cracking light cycle oil heavy fraction enters the hydrogenation reactor 420 through a heavy fraction pipeline 413 for circulation, and the catalytic cracking light cycle oil light fraction enters the third riser reactor 404 through a light fraction pipeline 418 and a fifth nozzle 419.

The invention will be further illustrated by the following examples, but is not to be construed as being limited thereto.

In the following examples and comparative examples, the hydrotreating catalyst loaded in the hydrogenation reactor was designated by the trade designation RN-32V, and the protectant was designated by the trade designation RG-1, both of which were manufactured by China petrochemical catalyst division. The filling volume ratio of the hydrotreating catalyst to the protective agent is 95: 5.

the catalytic cracking catalyst used in the riser reactor was commercially available as HAC, produced by china petrochemical catalyst division, and had physicochemical properties shown in table 1.

The heavy feedstock used was a blend of 90 wt% straight run wax oil and 10 wt% vacuum residue, the properties of which are given in table 2.

The light cycle oil circulation weight ratio is the weight of the recycled light cycle oil/the weight of the heavy raw material;

the recycle weight ratio of light ends is equal to the weight of light ends recycled/weight of heavy feedstock;

the recycle weight ratio of the heavy fraction is the weight of the heavy fraction of the hydrogenation recycle/the weight of the heavy feed;

hydrogen consumption is the fresh hydrogen consumption of the hydrogenation reactor per weight of fresh feedstock to the hydrogenation reactor.

The Research Octane Number (RON) of the obtained gasoline product is measured by a GB/T5487-.

The micro-reverse activity (MAT) of the regenerated catalyst is measured by a standard method of RIPP 92-90 (see "analytical methods of petrochemical industry (RIPP test method)", eds of Yangroi et al, science publishers, first edition of 9 months 1990, pp 263-268), and specific measurement conditions are as follows: catalyst: 5.0 g (20-40 mesh). Oil inlet amount: 1.56 g, reaction time: 70 seconds, reaction temperature: 460 ℃, agent/oil: 3.2, space velocity: 16 hours^-1。

Example 1

In this example, the process flow shown in fig. 1 is followed, the catalytic cracking light cycle oil is cut into a light fraction and a heavy fraction, the cutting temperature is 250 ℃, the heavy fraction is hydrogenated to obtain a hydrogenated heavy fraction, the heavy feedstock, the hydrogenated heavy fraction, and the light fraction are respectively injected into the first riser reactor through the first nozzle, the second nozzle, and the third nozzle, and the first nozzle, the third nozzle, and the second nozzle are arranged from bottom to top, that is, the heavy feedstock is first introduced into the first riser reactor, the light fraction is then introduced into the first riser reactor, and the hydrogenated heavy fraction is finally introduced into the first riser reactor.

The reaction conditions for the hydrotreatment were as follows: hydrogen partial pressure 8.0 MPa, average bed reaction temperature 360 deg.c and volume space velocity 0.5 hr^-1Hydrogen to oil volume ratio 1100Nm³/m³. The initial boiling point of the hydrogenated product (i.e., hydrogenated cycle oil) was 250 ℃, the bicyclic aromatic content was 19 wt%, and the hydrogen content was 11 wt%.

The catalytic cracking reaction conditions were as follows: the residence time of the reaction oil gas in the first riser reactor between the first nozzle and the second nozzle is 1 second, and the residence time of the reaction oil gas in the first riser reactor between the first nozzle and the third nozzle is 0.5 second. In unit time, the circulating weight ratio of the first catalytic cracking catalyst to the second catalytic cracking catalyst is 1: 0.05 and the second catalytic cracking catalyst is introduced between the first and second nozzles. The recycle weight ratio of the light fraction was 0.09, and the recycle weight ratio of the heavy fraction was 0.21. The reaction temperature is 550 ℃, the absolute pressure is 0.25 MPa, the weight ratio of the total weight of the first catalytic cracking catalyst and the second catalytic cracking catalyst to the heavy raw material is 8.1, the oil-gas retention time of the heavy raw material is 2.8 seconds, the weight ratio of the steam to the heavy raw material is 0.05, the ratio of the total weight of the first catalytic cracking catalyst and the second catalytic cracking catalyst to the total weight of the hydrogenated heavy fraction and the light fraction is 27, the ratio of the weight of the steam to the total weight of the hydrogenated heavy fraction and the light fraction is 0.01, and the micro-reverse activity of the first catalytic cracking catalyst and the second catalytic cracking catalyst is 64.

The reaction product distribution, hydrogen consumption and gasoline octane number of example 1 are shown in Table 4.

Comparative example 1

The process was substantially the same as that of example 1 except that the heavy feedstock, the hydrogenated heavy fraction and the light fraction were injected into the first riser reactor from the second nozzle, the third nozzle and the first nozzle, respectively, i.e., the light fraction was introduced into the first riser reactor, the hydrogenated heavy fraction was then introduced into the first riser reactor, and finally the heavy feedstock was introduced into the first riser reactor, the specific hydrogenation conditions and other catalytic cracking reaction conditions were the same as in example 1, and further, a second catalytic cracking catalyst was introduced into the first riser reactor above the second nozzle.

The reaction product distribution, hydrogen consumption and gasoline octane number of comparative example 1 are shown in Table 4.

Comparative example 2

The process was substantially the same as in example 1 except that the catalytically cracked light cycle oil was directly hydrotreated without cutting to obtain a hydrotreated cycle oil. The heavy feedstock and the hydrocyclized oil are injected into the first riser reactor via a first nozzle and a second nozzle, respectively. The reaction conditions for the catalytic cracking reaction are shown in table 3, and the residence time of oil and gas on the riser reactor between the second nozzle and the first nozzle is 0.2 seconds.

The reaction conditions for the hydrotreatment were as follows: hydrogen partial pressure 8.0 MPa, average bed reaction temperature 360 deg.c and volume space velocity 0.5 hr^-1Hydrogen to oil volume ratio 1100Nm³/m³. The initial boiling point of the hydrogenated product (i.e., hydrogenated cycle oil) was 170 ℃, the bicyclic aromatic content was 19 wt%, and the hydrogen content was 11 wt%.

The reaction product distribution, hydrogen consumption and gasoline octane number of comparative example 2 are shown in Table 4.

Comparative example 3

The process was essentially the same as comparative example 2, except that: the hydrogenated circulating oil and the heavy raw material are mixed and then are sprayed into the first riser reactor through the first nozzle. The reaction conditions of the catalytic cracking reaction are shown in Table 3, wherein the "weight ratio of catalyst to oil" is calculated based on the total amount of the mixed feed (including the heavy feedstock and the hydrogenated cycle oil), and the catalyst is all fed from the bottom of the first riser reactor.

The reaction conditions of the hydrotreatment were the same as in comparative example 2, and the initial boiling point of the hydrogenated cycle oil was 170 ℃, the bicyclic aromatic hydrocarbon content was 19 wt%, and the hydrogen content was 11 wt%.

The reaction product distribution, hydrogen consumption and gasoline octane number of comparative example 3 are shown in Table 4.

Comparative example 4

The process was essentially the same as comparative example 2, except that: all regenerated catalyst recycled back to the first riser reactor is fed from the bottom of the first riser reactor. The reaction conditions for the catalytic cracking reaction are shown in table 3.

The reaction conditions for the hydrotreatment were the same as in example 1, and the initial boiling point of the hydrogenated product (i.e., the hydrogenated cycle oil) was 170 ℃, the bicyclic aromatic hydrocarbon content was 19 wt%, and the hydrogen content was 11 wt%.

The reaction product distribution, hydrogen consumption and gasoline octane number of this comparative example are shown in Table 4.

Example 2

This example was carried out according to the process scheme shown in FIG. 2, except that the second riser reactor and the third riser reactor were both equal-diameter risers, the heavy feedstock was injected into the second riser reactor via the sixth nozzle, and the catalytically cracked light cycle oil was cut into light and heavy fractions at a temperature of 250 ℃. And (3) hydrogenating the heavy fraction to obtain hydrogenated heavy fraction, and spraying the hydrogenated heavy fraction and the hydrogenated light fraction into the third riser reactor through a fourth nozzle and a fifth nozzle respectively.

The reaction conditions in the second riser reactor were as follows: the reaction temperature is 500 ℃, the weight ratio of the catalyst to the oil is 6.1, the oil gas residence time is 2.8 seconds, the absolute pressure is 0.25 MPa, the weight ratio of the water vapor to the heavy raw material is 0.05, and the micro-reaction activity of the first catalytic cracking catalyst is 64.

The reaction conditions in the third riser reactor were as follows: the reaction temperature is 550 ℃, the absolute pressure is 0.25 MPa, the weight ratio of the weight of the second catalytic cracking catalyst to the hydrogenated heavy fraction is 40, the oil-gas retention time of the hydrogenated heavy fraction is 2.8 seconds, the weight ratio of the water vapor to the hydrogenated cycle oil is 0.01, the weight ratio of the weight of the second catalytic cracking catalyst to the light fraction is 115, the weight ratio of the water vapor to the light fraction is 0.02, and the micro-reaction activity of the second catalytic cracking catalyst is 64. The oil gas residence time of the third riser reactor between the fifth nozzle and the fourth nozzle was 0.5 seconds, the circulating weight ratio of the light fraction was 0.09, and the circulating weight ratio of the heavy fraction was 0.25.

The hydrotreating conditions were the same as in example 1, the initial boiling point of the hydrogenated heavy fraction being 170 ℃, the bicyclic aromatic content being 19% by weight, and the hydrogen content being 11% by weight.

The reaction product distribution, hydrogen consumption and gasoline octane number of this example are shown in Table 6.

Example 3

The method is basically the same as that of the embodiment 2, except that the second riser reactor is a reducing riser reactor, the second riser reactor is coaxially provided with a first reaction section and a second reaction section from bottom to top, the inner diameter of the second reaction section is larger than that of the first reaction section, the top opening of the third riser reactor is communicated with the second reaction section through a horizontal pipe, and the reaction conditions of the reaction sections are as follows: the temperature is 500 ℃, the absolute pressure is 0.25 MPa, the micro-reverse activity of the regenerated catalyst is 64, the weight ratio of the catalyst to the oil is 6.1, the oil gas retention time is 1.2 seconds, and the weight ratio of the water vapor to the heavy raw material is 0.06; the reaction conditions of the expanding section are as follows: the temperature was 490 deg.C and the oil gas residence time was 5 seconds. The hydrotreating conditions and other catalytic cracking conditions were the same as in example 2.

Comparative example 5

The procedure is as in example 2, except that: the conventional arrangement of a second riser reactor and a third riser reactor is adopted, namely the second riser reactor and the third riser reactor are arranged in parallel, the top outlets of the second riser reactor and the third riser reactor are directly communicated with a settler, and reaction materials taken out from the top of the third riser reactor are directly sent into the settler; and the hydrogenated heavy fraction and the hydrogenated light fraction are sprayed into the third riser reactor through a fourth nozzle and a fifth nozzle.

The reaction conditions for hydrotreating and catalytic cracking were the same as in example 2. The initial boiling point of the hydrogenated product (i.e., hydrogenated cycle oil) was 170 ℃, the bicyclic aromatic content was 19 wt%, and the hydrogen content was 11 wt%.

The reaction product distribution, hydrogen consumption and gasoline octane number of this comparative example are shown in Table 6.

Comparative example 6

By adopting the reaction system of example 2, the second/third riser reactors used in this comparative example were all equal-diameter risers, the heavy feedstock was injected into the second riser reactor via the sixth nozzle, all of the catalytic cracking cycle oil obtained by separation was hydrotreated, and the hydrogenated cycle oil obtained was injected into the third riser reactor via the fifth nozzle. The reaction conditions for the catalytic cracking reaction are shown in Table 5.

The reaction conditions for the hydrotreatment were as follows: hydrogen partial pressure 8.0 MPa, average bed reaction temperature 360 deg.c and volume space velocity 0.5 hr^-1Hydrogen to oil volume ratio 1100Nm³/m³. The initial boiling point of the hydrogenated cycle oil is 170 ℃, the content of the bicyclic aromatic hydrocarbon is 19 weight percent, and the content of hydrogen is 11 weight percent.

Comparative example 7

The operation was carried out as described in comparative example 6, with the difference that: the second riser reactor is a reducing riser reactor, and is provided with a pre-lifting section, a first reaction section I, a second reaction section II and an outlet section from bottom to top in sequence, wherein the reaction conditions of the reaction sections are as follows: the temperature is 500 ℃, the absolute pressure is 0.25 MPa, the micro-reverse activity of the regenerated catalyst is 64, the weight ratio of the catalyst to the oil is 6.1, the oil gas retention time is 1.2 seconds, and the weight ratio of the water vapor to the heavy raw material is 0.06; the reaction conditions of the expanding section are as follows: the temperature was 490 deg.C and the oil gas residence time was 5 seconds.

The reaction conditions in the third riser reactor and the hydrotreating reaction conditions were the same as in comparative example 6, and the circulating weight ratio of the light cycle oil was 0.07. The initial boiling point of the hydrogenated product (i.e., hydrogenated cycle oil) was 170 ℃, the bicyclic aromatic content was 19 wt%, and the hydrogen content was 11 wt%.

Comparative example 8

The operation was carried out as described in comparative example 6, with the difference that: only a second riser reactor is arranged, and a third riser reactor is not arranged; hydrogenated cycle oil obtained by hydrogenating the catalytic cracking light cycle oil and the heavy raw material are mixed and sprayed into the second riser reactor through the sixth nozzle. The catalytic cracking reaction conditions for the second riser reactor are shown in table 5.

The reaction conditions for the hydrotreatment were the same as in comparative example 6. The initial boiling point of the hydrogenated cycle oil is 170 ℃, the content of the bicyclic aromatic hydrocarbon is 19 weight percent, and the content of hydrogen is 11 weight percent.

Comparative example 9

The operation was carried out as described in comparative example 6, with the difference that: the conventional arrangement of a second riser reactor and a third riser reactor is adopted, namely the second riser reactor and the third riser reactor are arranged in parallel, the top outlets of the second riser reactor and the third riser reactor are directly communicated with a settler, and reaction materials taken out from the top of the third riser reactor are directly sent into the settler; and the hydrogenated light cycle oil after the hydrogenation of the catalytic cracking light cycle oil is sprayed into the third riser reactor through a fifth nozzle. The reaction conditions for the catalytic cracking reaction are shown in Table 5.

The reaction conditions for the hydrotreatment were the same as in comparative example 6. Namely, the initial boiling point of the hydrogenated cycle oil is 170 ℃, the content of the bicyclic aromatic hydrocarbon is 19 weight percent, and the content of hydrogen is 11 weight percent.

It can be seen from tables 4 and 6 that the process of the present invention has a higher yield of high octane gasoline.

The preferred embodiments of the present invention have been described in detail, however, the present invention is not limited to the specific details of the above embodiments, and various simple modifications may be made to the technical solution of the present invention within the technical idea of the present invention, and these simple modifications are within the protective scope of the present invention.

It should be noted that the various technical features described in the above embodiments can be combined in any suitable manner without contradiction, and the invention is not described in any way for the possible combinations in order to avoid unnecessary repetition.

In addition, any combination of the various embodiments of the present invention is also possible, and the same should be considered as the content of the present invention as long as it does not depart from the gist of the present invention.

TABLE 1

TABLE 2

Raw oil name	Mixing the raw materials
		Density (20 ℃) kg/m³	916.8
Freezing point, DEG C	32
		Refractive index (70 ℃ C.)	1.4968
Carbon residue, by weight%	2.67
		Average molecular weight	404
Distillation range, deg.C
		Initial boiling point	294
5% by weight	361
		10% by weight	381
30% by weight	422
		50% by weight	451
70% by weight	497
		Sulfur content, wt.%	1.1
Nitrogen content, wt.%	0.24
		Hydrogen content, wt.%	12.6
Metal content, mg/kg
		Ni	6.6
V	1.2

TABLE 3

TABLE 4

TABLE 5

TABLE 6

Claims

1. A process for catalytic cracking, the process comprising:

wherein, the catalytic cracking reaction in the step a and the catalytic cracking reaction in the step e are carried out in a first riser reactor, the heavy raw material and the hydrogenated heavy fraction are respectively sprayed into the first riser reactor from a first nozzle and a second nozzle which are arranged from bottom to top at intervals and react from bottom to top, the light fraction is sprayed into the first riser reactor from a third nozzle, the third nozzle is positioned between the first nozzle and the second nozzle along the height direction of the first riser reactor, and the introduction position of the first catalytic cracking catalyst is at the bottom of the first riser reactor, or

The catalytic cracking reaction in the step a and the catalytic cracking reaction in the step e are respectively carried out in a second riser reactor and a third riser reactor, the top opening of the third riser reactor is communicated with the middle upper part of the second riser reactor through a horizontal pipe, and the second reaction product and the second spent catalyst are fed into the middle upper part of the second riser reactor from the top opening of the third riser reactor and are discharged from the top opening of the second riser reactor together with the first reaction product and the first spent catalyst; and the hydrogenated heavy fraction is sprayed into the third riser reactor from a fourth nozzle, the light fraction is sprayed into the third riser reactor from a fifth nozzle, and the fourth nozzle is positioned above the fifth nozzle along the height direction of the third riser reactor.

2. The process of claim 1 wherein the second catalytic cracking catalyst is introduced at a location between the first and third nozzles.

3. The method of claim 2, wherein the residence time of the reaction oil gas in the first riser reactor between the first nozzle and the second nozzle is between 0.05 and 3 seconds; in the first riser reactor between the first nozzle and the third nozzle, the residence time of the reaction oil gas is 0.01-2 seconds;

4. The process of claim 1 wherein the first riser reactor has a first catalytic cracking catalyst and a second catalytic cracking catalyst in a circulating weight ratio per unit time of 1: (0.02-1), the circulating weight ratio of the light fraction is 0.01-1, and the circulating weight ratio of the heavy fraction is 0.01-1;

5. The method of claim 1 wherein the reaction conditions of the first riser reactor comprise: the reaction temperature is 520 ℃ and 650 ℃, the absolute pressure is 0.15-0.4 MPa, the weight ratio of the total weight of the first catalytic cracking catalyst and the second catalytic cracking catalyst to the heavy raw material is 1-50, the oil gas retention time of the heavy raw material is 1-10 seconds, and the weight ratio of the water vapor to the heavy raw material is 0.01-0.5; the proportion of the total weight of the first catalytic cracking catalyst and the second catalytic cracking catalyst to the total weight of the hydrogenated heavy fraction and the light fraction is 5-100, the proportion of the weight of the steam to the total weight of the hydrogenated heavy fraction and the light fraction is 0.01-0.3, and the micro-reverse activities of the first catalytic cracking catalyst and the second catalytic cracking catalyst are not less than 60;

6. The method of claim 1 wherein the first, second, and third riser reactors are each a constant diameter riser reactor or a variable diameter riser reactor.

7. The method according to claim 1, wherein the second riser reactor is coaxially provided with a first reaction section and a second reaction section from bottom to top, the inner diameter of the second reaction section is larger than that of the first reaction section, and the top opening of the third riser reactor is communicated with the second reaction section through a horizontal pipe.

8. The process of claim 1, wherein the heavy feedstock is selected from one or more of straight run wax oil, coker wax oil, deasphalted oil, hydrofinished oil, hydrocracked tail oil, vacuum resid, and atmospheric resid.

9. The process of claim 1, wherein the first and second catalytic cracking catalysts each comprise 10-50 wt% of a zeolite selected from one or more of rare earth-containing or non-containing Y-type zeolite, HY-type zeolite, USY-type zeolite, and Beta zeolite, 5-90 wt% of an inorganic oxide selected from silica and/or aluminum trioxide, and 0-70 wt% of a clay selected from kaolin and/or halloysite.

10. The process of claim 1, wherein the hydrotreating catalyst comprises an active metal component which is a group VIB metal and/or a non-noble group VIII metal, and a support which is at least one selected from alumina, silica and amorphous silica-alumina.

11. The method of claim 10, wherein the active metal component is nickel-tungsten, nickel-tungsten-cobalt, nickel-molybdenum, or cobalt-molybdenum.

12. The method of claim 1, wherein the hydrotreating conditions comprise: hydrogen partial pressure of 5.0-22.0 MPa, reaction temperature of 330-^-1Hydrogen oil volume ratio of 100-2000Nm³/m³。

13. The process of claim 1, wherein the hydrogenated heavy fraction has a bicyclic aromatic content of no greater than 20 wt% and a hydrogen content of no less than 10 wt%.

14. A system suitable for use in the catalytic cracking process of any of claims 1-13, the system comprising a reaction unit, a regenerator, a hydrogenation reactor, a settler, and a product separation unit;

wherein the reaction unit comprises a first riser reactor, the first riser reactor is provided with a first catalytic cracking catalyst inlet positioned at the bottom, a second catalytic cracking catalyst inlet positioned at the middle part, a first nozzle used for feeding heavy raw materials, a second nozzle used for feeding hydrogenated heavy fractions, a third nozzle used for feeding light fractions and a top opening used for feeding oil, the first nozzle is positioned below the second nozzle and along the height direction of the first riser reactor, the third nozzle is positioned between the first nozzle and the second nozzle, the second nozzle is communicated with a hydrogenated heavy fraction outlet of the hydrogenation reactor, the third nozzle is communicated with a light fraction outlet of the product separation device, the top opening of the first riser reactor is communicated with the settler, and a catalyst outlet of the settler is communicated with a catalyst inlet of the regenerator, the oil gas outlet of the settler is communicated with the raw material inlet of the product separation device, and the catalyst outlet of the regenerator is communicated with the first catalytic cracking catalyst inlet and the second catalytic cracking catalyst inlet of the first riser reactor; or

The reaction unit comprises a second riser reactor and a third riser reactor, the second riser reactor is provided with a heavy raw material inlet at the lower part, a catalytic cracking catalyst inlet at the bottom and a top opening for sending out oil, the third riser reactor is provided with a fourth nozzle for sending in hydrogenated heavy fraction, a fifth nozzle for sending in light fraction, a catalytic cracking catalyst inlet at the bottom and a product outlet at the top, and the fourth nozzle is positioned above the fifth nozzle along the height direction of the third riser reactor; the fourth nozzle is communicated with a hydrogenation heavy fraction outlet of the hydrogenation reactor, the fifth nozzle is communicated with a light fraction outlet of the product separation device, a top opening of the third riser reactor is communicated with the middle upper part of the second riser reactor through a horizontal pipe, a top opening of the second riser reactor is communicated with the settler, a catalyst outlet of the settler is communicated with a catalyst inlet of the regenerator, an oil gas outlet of the settler is communicated with a raw material inlet of the product separation device, and a catalyst outlet of the regenerator is communicated with catalytic cracking catalyst inlets of the second riser reactor and the third riser reactor.

15. The system of claim 14, wherein the second catalytic cracking catalyst inlet is located between the first nozzle and a third nozzle.

16. The system of claim 14 wherein the first, second, and third riser reactors are each a constant diameter riser reactor or a variable diameter riser reactor.

17. The system of claim 14, wherein the second riser reactor is coaxially provided with a first reaction section and a second reaction section from bottom to top, the inner diameter of the second reaction section is larger than that of the first reaction section, and the top opening of the third riser reactor is communicated with the second reaction section through a horizontal pipe.