CN112745901B

CN112745901B - Catalytic conversion method and catalytic conversion device for producing low-carbon olefin

Info

Publication number: CN112745901B
Application number: CN201911046129.5A
Authority: CN
Inventors: 马文明; 朱根权; 杨超
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: 2019-10-30
Filing date: 2019-10-30
Publication date: 2023-05-05
Anticipated expiration: 2039-10-30
Also published as: CN112745901A

Abstract

The invention provides a catalytic conversion method and a catalytic conversion device for producing low-carbon olefin. The method comprises the following steps: contacting a first hydrocarbon feedstock with a cracking catalyst in a first riser reactor to produce a first oil mixture; contacting the second hydrocarbon feedstock with a cracking catalyst in a second riser reactor to produce a second oil mixture; and mixing the first oil mixture and the second oil mixture, then entering a third reactor for reaction, and converting the cracking catalyst into a spent catalyst after reaction and entering a stripper, wherein part of the spent catalyst is introduced into the second riser reactor. The catalytic conversion method and device of the invention effectively control the properties of the catalyst in the second riser reactor and the third reactor by introducing the catalyst to be regenerated in the second riser reactor, thereby improving the yield of the low-carbon olefin.

Description

Catalytic conversion method and catalytic conversion device for producing low-carbon olefin

Technical Field

The invention relates to the field of petroleum processing, in particular to a catalytic conversion method and a catalytic conversion device for producing low-carbon olefin.

Background

The low-carbon olefin such as propylene, butene, etc. is a basic chemical raw material, and is mainly derived from steam thermal cracking of MTO, butane, LPG, condensate, naphtha, hydrocracking tail oil, gas oil, etc. and catalytic cracking of vacuum fractions at present. With the adoption of new light raw materials for steam cracking, the distribution of products is changed, for example, ethane is adopted as the steam cracking raw materials, the proportion of ethylene in the products is obviously improved compared with naphtha, and the yields of propylene and butene are reduced. The catalytic cracking process can produce more low-carbon olefin, and is an effective supplementary measure for preparing ethylene by steam thermal cracking. However, the conventional catalytic cracking process has a low yield of light olefins, which is not more than 15% of the feedstock, and it is difficult to meet the market demand, so it is very necessary to develop a catalytic cracking technology capable of treating heavy feedstock and producing high yields of light olefins.

A method for the extensive use of shape selective cracking aids in the catalytic cracking of heavy feedstocks is disclosed in US 5997728. The auxiliary agent consists of an amorphous matrix and ZSM-5 zeolite added into the amorphous matrix, wherein the system reserve is at least 10%, so that the proportion of ZSM-5 in the catalyst exceeds 3%. The method can greatly improve propylene and butylene without additionally increasing the aromatic hydrocarbon yield and losing the gasoline yield.

CN1031834a discloses a catalytic conversion process for producing light olefins. The method takes petroleum fractions, residual oil or crude oil with different boiling ranges as raw materials, takes a mixture containing Y zeolite and pentasil zeolite as a catalyst, adopts a fluidized bed or a moving bed as a reactor, and has the following reaction conditions: the temperature is 500-650 ℃, the pressure is 0.15-0.30 MPa, and the weight hourly space velocity is 0.2-20 hours ^-1 Catalyst oil ratio is 2-12, and the catalyst after reaction returns to the reactor for recycling after being burnt and regenerated. The present process is capable of producing more propylene and butene than conventional catalytic cracking and steam cracking.

A catalytic cracking process for producing propylene is disclosed in CN102690683 a. The method adopts a double-riser configuration, wherein the first riser reactor is used for treating heavy hydrocarbon oil, a catalyst containing Y-type zeolite and beta-type zeolite is used, the second riser reactor is used for treating light hydrocarbon, and a shape selective zeolite with the pore diameter smaller than 0.7nm is used. The method adopts two different catalysts, and divides the stripping zone and the regeneration zone into two independent parts through the partition plates respectively, thereby increasing the complexity of the device and being unfavorable for operation.

CN102206509a discloses a hydrocarbon catalytic conversion process for producing propylene and light aromatic hydrocarbons. The method adopts a combined reactor form of double lifting pipes and a fluidized bed reactor, wherein heavy hydrocarbon and a cracking catalyst containing modified beta zeolite are in contact reaction in a first reactor, C4 hydrocarbon fraction and/or light gasoline fraction and the cracking catalyst containing modified beta zeolite are introduced into a third reactor for continuous reaction after being in contact reaction in a second reactor, and the third reactor is the fluidized bed reactor, thereby creating conditions for secondary cracking reaction of the gasoline fraction and further improving the yield of propylene and light aromatic hydrocarbon.

CN103131464a discloses a hydrocarbon catalytic conversion process for producing propylene and light aromatic hydrocarbons. The method comprises the steps of carrying out contact reaction on petroleum hydrocarbon and a catalytic cracking catalyst in a riser, enabling reaction effluent to enter a fluidized bed reactor without separation, carrying out oligomerization, cracking and aromatization reactions by contacting the reaction effluent with the catalyst which is introduced and is subjected to pore channel modification treatment, separating to obtain products comprising low-carbon olefin and light aromatic hydrocarbon, and separating the products into two parts after stripping and regeneration, wherein one part of the carbon deposition catalyst is recycled in the riser, and the other part of the carbon deposition catalyst is firstly sent to a catalyst pore channel modification area, then is contacted and reacted with a contact agent and is sent to a fluidized bed for use. The method has higher heavy oil conversion capability and high propylene selectivity to heavy hydrocarbon raw materials.

The above techniques promote the conversion of heavy hydrocarbon feedstock and increase the selectivity of light olefins by adjusting the catalyst formulation and employing a combined reactor form of riser in combination with fluidized bed, but do not involve flexible control of the reaction environment in the fluidized bed reactor, and the yield of light olefins remains to be further improved.

It is noted that the information disclosed in the foregoing background section is only for enhancement of understanding of the background of the invention and therefore it may contain information that does not form the prior art that is already known to a person of ordinary skill in the art.

Disclosure of Invention

The invention aims to provide a catalytic conversion method and a catalytic conversion device for producing low-carbon olefin, which can effectively regulate the reaction environment, so as to improve the yield of the low-carbon olefin.

In order to achieve the above purpose, the present invention adopts the following technical scheme:

a catalytic conversion process for producing light olefins comprising:

contacting a first hydrocarbon feedstock with a cracking catalyst in a first riser reactor to produce a first oil mixture;

contacting the second hydrocarbon feedstock with a cracking catalyst in a second riser reactor to produce a second oil mixture; and

and mixing the first oil mixture and the second oil mixture, then entering a third reactor for reaction, and converting the cracking catalyst into a spent catalyst after reaction and entering a stripper, wherein part of the spent catalyst is introduced into the second riser reactor.

In some embodiments, the reaction conditions of the first riser reactor include: the reaction temperature is 480-620 ℃, preferably 520-600 ℃; the ratio of the agent to the oil is 2-25, preferably 3-20; the reaction time is 1 to 15 seconds, preferably 2 to 10 seconds.

In some embodiments, the reaction conditions of the second riser reactor include: the reaction temperature is 560-720 ℃, preferably 580-680 ℃, the catalyst-to-oil ratio is 3-40, preferably 5-30, and the reaction time is 0.5-10 seconds, preferably 1-5 seconds.

In some embodiments, the reaction temperature of the third reactor is 520 to 700 ℃ and the weight hourly space velocity is 1 to 30 hours ^-1 The catalyst density is 100-500 kg/m ³ The height of the bed layer is 1/2-4/5 of the height of the bed layer reactor, and the pressure in the reactor is 0.1-0.4 MPa.

In some embodiments, the first hydrocarbon feedstock is selected from one or more of vacuum wax oil, atmospheric residuum, vacuum residuum, coker wax oil, deasphalted oil, furfural extract oil, coal liquefaction oil, oil sand oil, shale oil, distillate oil obtained by F-T synthesis, or animal and vegetable oil.

In some embodiments, the second hydrocarbon feedstock is a mixture of C4-C8 hydrocarbons, the mixture of C4-C8 hydrocarbons having an olefin content of greater than 50wt%, preferably greater than 60wt%.

In some embodiments, the cracking catalyst comprises a cracking active component, clay, and a binder, wherein the cracking active component comprises a molecular sieve having an MFI structure,

Wherein the content of the cracking active component is 20 to 70wt percent, preferably 30 to 50wt percent, based on the dry weight of the cracking catalyst; the clay content is 15-60 wt%, preferably 30-50 wt%; the binder content is 20 to 35wt%, preferably 20 to 30wt%.

In some embodiments of the present invention, in some embodiments,

the cracking active component further comprises an optional Y molecular sieve and an optional beta molecular sieve, wherein the content of the Y molecular sieve is 0-90 wt%, preferably 50-80 wt%, based on the total weight of the cracking active component; the content of the molecular sieve with the MFI structure is 1 to 50wt%, preferably 10 to 40wt%, and the content of the beta molecular sieve is 0 to 50wt%, preferably 10 to 40wt%.

In some embodiments, further comprising introducing a lift gas to a middle portion of the second riser reactor and the lift gas addition port is located below the spent catalyst addition port, wherein the lift gas is selected from one or more of steam, nitrogen, dry gas, preferably steam.

In some embodiments, further comprising introducing a lift gas to the third reactor, preferably introducing a lift gas to the bottom of the third reactor, wherein the lift gas is selected from one or more of steam, nitrogen, dry gas, preferably steam.

In some embodiments, the method further comprises introducing the remaining spent catalyst into a regenerator for regeneration to obtain regenerated catalyst, and returning the regenerated catalyst to the first riser reactor and the second riser reactor respectively.

In another aspect, the present invention also provides a catalytic conversion apparatus for producing light olefins, comprising:

a combined reactor, wherein the combined reactor comprises a first riser reactor, a second riser reactor and a third reactor, the first riser reactor and the second riser reactor being connected to the third reactor, respectively;

a stripper located below the third reactor and in communication with the third reactor;

the regenerator is connected with the stripper and is respectively connected with the first riser reactor and the second riser reactor through regenerated catalyst pipelines; and

and the spent catalyst conveying pipe is respectively communicated with the stripper and the second riser reactor.

In some embodiments, the second riser reactor is provided with a riser gas distribution ring in the middle.

In some embodiments, the lift gas distribution ring is disposed below the junction of the spent catalyst transfer pipe and the second riser reactor.

In some embodiments, the bottom of the third reactor is provided with a lift gas distribution ring.

In some embodiments, a settler is also included, the settler in communication with the third reactor.

In some embodiments, a separation device is also included, the separation device being disposed in an upper portion of the settler.

In some embodiments, the first riser reactor and the second riser reactor are each independently selected from one or more combinations of an equal diameter riser reactor, an equal linear velocity riser reactor, and a variable diameter riser reactor, and the third reactor is selected from one or more combinations of a fixed fluidized bed reactor, a bulk fluidized bed reactor, a bubbling bed reactor, a turbulent bed reactor, a fast bed reactor, a entrained bed reactor, and a dense phase fluidized bed reactor.

The catalytic conversion method and the catalytic conversion device of the invention effectively control the properties of the catalyst in the second riser reactor and the third reactor by introducing the spent catalyst in the second riser reactor, thereby improving the yield of the low-carbon olefin.

In addition, the catalytic conversion method and the catalytic conversion device can effectively regulate the flow of the lifting gas through the additionally arranged lifting gas distribution ring, flexibly control the reaction environment of the reactor, further promote the secondary cracking of the gasoline fraction and improve the yield of the low-carbon olefin in the product.

Drawings

Fig. 1 is a schematic structural view of a catalytic converter according to an embodiment of the present invention;

FIG. 2 is a schematic cross-sectional view of the lift gas distribution ring of FIG. 1;

wherein reference numerals are as follows:

1-first reactor

11-first hydrocarbon feedstock line

12-first pre-lift gas line

13-cycle oil pipeline

2-second reactor

21-second hydrocarbon feed line

22-second pre-lift gas line

23-second hydrocarbon feedstock line

24-Lift gas pipeline

25-lift gas distribution ring

3-third reactor

31-Lift gas pipeline

32-lift gas distribution ring

321-lift gas inlet line

322-lifting gas distribution port

4-settler

41. 42-cyclone separator

43-plenum chamber

44-cracked oil gas

5-stripper

51. 52-spent catalyst transfer tube

53-stripping gas line

54-stripping gas distribution ring

55-stripping baffle

6-scorching tank

7-regenerator

71-second regenerated catalyst line

72-first regenerated catalyst line

73. 74-cyclone separator

75-plenum chamber

76-regeneration flue gas outlet

Detailed Description

The technical scheme of the invention is further described below according to specific embodiments. The scope of the invention is not limited to the following examples, which are given for illustrative purposes only and do not limit the invention in any way.

In the present invention, any matters or matters not mentioned are directly applicable to those known in the art without modification except for those explicitly stated. Moreover, any embodiment described herein can be freely combined with one or more other embodiments described herein, and the technical solutions or ideas thus formed are all considered as part of the original disclosure or original description of the present invention, and should not be considered as new matters not disclosed or contemplated herein unless such combination would obviously be unreasonable to one skilled in the art.

All of the features disclosed in this invention may be combined in any combination which is understood to be disclosed or described in this invention unless the combination is obviously unreasonable by those skilled in the art. 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.

According to one embodiment of the present invention, a catalytic conversion process for producing light olefins comprises:

the first oil agent mixture and the second oil agent mixture are mixed and then enter a third reactor for reaction, the reacted oil agent mixture is separated by a separating device, the cracking catalyst is converted into a spent catalyst after reaction and enters a stripper,

wherein a portion of the spent catalyst is introduced into the second riser reactor during the catalytic conversion process.

In the catalytic conversion method, oil gas separated from the oil-gas mixture is separated by a subsequent product separation system to obtain products such as dry gas, cracked gas, gasoline, light oil, slurry oil and the like. The cracked gas can be separated and refined to obtain a polymerization grade propylene product and a C4 fraction rich in olefin. The gasoline is firstly cut into light and medium gasoline fraction segments, and part or all of C4 fraction rich in olefin is returned to the second riser reactor for reaction, and part or all of light gasoline is returned to the second riser reactor for reaction.

The oil gas is rapidly separated from the reacted carbon deposition catalyst through the separation device, so that the yield of dry gas can be reduced, the conversion of propylene and butene after generation is inhibited, and the product separation system used later can be various separation systems in the prior art.

In the catalytic conversion method of the invention, the first hydrocarbon raw material is preheated to 180-340 ℃ and then enters the first riser reactor, and the reaction temperature with the cracking catalyst is 480-620 ℃, preferably 520-600 ℃; the catalyst to oil ratio (weight ratio of cracking catalyst to first hydrocarbon feedstock) is 2 to 25, preferably 3 to 20; the reaction time is 1 to 15 seconds, preferably 2 to 10 seconds; the pressure (absolute pressure) in the reactor is 0.1-0.4 MPa, preferably 0.15-0.35 MPa; the first oil mixture produced by the contact reaction under the condition that the atomized water vapor accounts for 10 to 30 weight percent, preferably 10 to 20 weight percent of the feeding amount of the heavy raw materials is introduced into the third reactor.

The first hydrocarbon raw material is selected from one or more of vacuum wax oil, normal pressure residual oil, vacuum residual oil, coking wax oil, deasphalted oil, furfural refined raffinate oil, coal liquefied oil, oil sand oil, shale oil, distillate oil obtained by F-T synthesis or animal and vegetable oil. The first hydrocarbon feedstock is subjected to a primary cracking reaction in a first reactor to convert from macromolecular reactants to small molecular products.

In the catalytic conversion method of the present invention, the second hydrocarbon raw material is preheated to 100-150 ℃ and then enters the second riser reactor, and the reaction temperature with the cracking catalyst is 560-720 ℃, preferably 580-680 ℃, the catalyst-oil ratio is 3-40, preferably 5-30, and the reaction time is 0.5-10 seconds, preferably 1-5 seconds; the pressure (absolute pressure) in the reactor is 0.1-0.4 MPa, preferably 0.15-0.35 MPa; the atomized water vapor preferably comprises 1 to 10 wt.%, preferably 2 to 5 wt.% of the weight of the feed, and the resulting second oil mixture is introduced into the third reactor.

The second hydrocarbon feedstock is a mixture of C4 to C8 hydrocarbons which is rich in olefins, preferably in an olefin content of greater than 50 wt%, more preferably greater than 60 wt%.

The mixture of C4 to C8 hydrocarbons comprises a C4 fraction and a light gasoline fraction produced by the catalytic conversion apparatus of the present invention (i.e., from the product separation system of the present invention) or by another apparatus: the C4 and light gasoline fractions produced by the other apparatus may be selected from one or more of a catalytically cracked C4 and light gasoline fraction, a coker light gasoline fraction, a visbreaker light gasoline fraction, and other light gasoline fractions produced by refinery or chemical processes, preferably the C4 and light gasoline fractions produced by the apparatus. The weight ratio of the mixture of C4-C8 hydrocarbons injected into the second reactor to the first hydrocarbon feedstock injected into the first riser reactor is from 0.05 to 0.20:1, preferably from 0.08 to 0.15:1.

In the catalytic conversion method of the present invention, after the first oil mixture from the first riser reactor and the second oil mixture from the second riser reactor enter the third reactor, the reaction is carried out under the conditions that the reaction temperature is 520-700 ℃, preferably 580-650 ℃, the weight hourly space velocity is 1-30 hours-1, preferably 5-20 hours-1, the catalyst density is 100-500 kg/m3, preferably 200-400 kg/m3, the bed height is 1/2-4/5 of the bed reactor height, preferably 1/2-3/4 of the bed reactor height, and the pressure in the reactor is 0.1-0.4 MPa (absolute pressure), preferably 0.15-0.3 MPa.

In the catalytic conversion method of the invention, the spent catalyst introduced into the middle part of the second riser reactor is usually spent catalyst from a stripper, and the flow rate of the spent catalyst can be regulated through the opening of a slide valve, so that the properties of the catalyst in the second riser reactor and the third reactor, including the activity, metal content, temperature and the like of the catalyst, are effectively controlled.

In the catalytic conversion process of the present invention, the cracking catalyst used in the first riser reactor and the second riser reactor are regenerated catalysts obtained by passing through a regenerator, the cracking catalyst comprising a cracking active component, clay and binder, wherein the cracking active component comprises a molecular sieve having an MFI structure, such as one or more selected from the group consisting of ZRP zeolite, phosphorous ZRP zeolite (CN 1194181A), rare earth ZRP zeolite (CN 1052290 a), phosphorous ZRP zeolite and rare earth ZRP zeolite (CN 1147420 a), phosphorous ZRP zeolite and alkaline earth ZRP zeolite (CN 1211470 a) and phosphorous ZRP zeolite and transition metal ZRP zeolite (CN 1465527 a), preferably phosphorous ZRP zeolite and rare earth ZRP zeolite; the clay is selected from various clays which can be used as catalyst components, such as kaolin, montmorillonite, bentonite, etc. The binder is selected from one or two or three of silica sol, alumina sol and pseudo-boehmite, wherein the preferred binder is a double-aluminum binder of alumina sol and pseudo-boehmite.

The content of the cracking active components is 20 to 70 weight percent, preferably 30 to 50 weight percent, based on the dry weight of the cracking catalyst; the clay content is 15-60 wt%, preferably 30-50 wt%; the content of the binder is 20 to 35wt%, preferably 20 to 30wt%.

The Y molecular sieve can be selected from one or more of HY, USY, REUSY, REY, REHY, DASY, REDASY or Y-type molecular sieves obtained by treating various metal oxides.

The beta molecular sieve is a beta molecular sieve modified by phosphorus and transition metal M, wherein M is selected from one or more of Fe, co, ni, cu, mn, zn and Sn. The beta molecular sieve modified by phosphorus and transition metal M can be prepared by various methods, for example, phosphorus and transition metal M can be introduced in the process of synthesizing the beta molecular sieve, or the steps of ammonium exchange, phosphorus modification, transition metal M modification, calcination treatment and the like are adopted after the synthesis of the beta molecular sieve to introduce phosphorus and transition metal M. Specific preparation of beta molecular sieves can be found in CN1035668C and CN1041616C.

In the catalytic conversion method, in order to further control the reaction environment of the third reactor, including bed height, bed density, bed space velocity and the like, lifting gas can be introduced into the middle part of the second riser reactor, wherein the inlet of the lifting gas is positioned below the inlet of the catalyst to be regenerated (according to the flow direction of the second hydrocarbon raw material, the inlet of the lifting gas is positioned before the inlet of the catalyst to be regenerated), so that the catalyst to be regenerated is driven to flow upwards by the lifting gas and finally enter the third reactor, and the lifting gas can be introduced into the bottom of the third reactor, and the two modes can be adopted simultaneously or respectively, so that the secondary cracking of gasoline fractions is further promoted, and the yield of low-carbon olefins in products is improved. The lifting gas is selected from one or more of steam, nitrogen and dry gas, preferably steam.

In the catalytic conversion method of the invention, the spent catalyst (the rest part after being introduced into the second riser reactor) can be introduced into the regenerator for regeneration to obtain regenerated catalyst, and the regenerated catalyst is respectively returned to the first riser reactor and the second riser reactor, thereby being circularly applied.

The regenerated catalyst is usually at a temperature higher than 700 ℃ due to the regeneration requirement, and for this purpose, the regenerated catalyst is subjected to heat-extracting and cooling treatment. The regenerated catalyst introduced into the first riser reactor is usually cooled to a temperature of 520 to 680 ℃, preferably 550 to 620 ℃, and the first hydrocarbon feedstock is contacted with the regenerated catalyst at the bottom of the first riser reactor and reacted, and the mixing temperature after the first hydrocarbon feedstock is contacted with the regenerated catalyst is 520 to 680 ℃, preferably 540 to 610 ℃. The regenerated catalyst introduced into the second riser reactor is usually cooled to a temperature of 550 to 680 ℃, preferably 580 to 660 ℃, and the second hydrocarbon feedstock is contacted with the regenerated catalyst at the bottom of the second riser reactor and reacted, and the mixing temperature after the contact of the second hydrocarbon feedstock with the regenerated catalyst is 540 to 680 ℃, preferably 550 to 630 ℃.

In another aspect, the present invention further provides a catalytic converter for producing light olefins, and fig. 1 is a schematic structural diagram of a catalytic converter according to an embodiment of the present invention, as shown in fig. 1, where the catalytic converter includes: a combined reactor, a stripper 5, a regenerator 7 and a spent catalyst transfer line 51.

In the catalytic conversion device of the invention, the combined reactor adopts a mode of combining double lifting pipes with a fluidized bed, and specifically comprises a first lifting pipe reactor 1, a second lifting pipe reactor 2 and a third lifting pipe reactor 3, wherein the first lifting pipe reactor 1 and the second lifting pipe reactor 2 are respectively connected to the third lifting pipe reactor 3, and the third lifting pipe reactor 3 is preferably positioned above the first lifting pipe reactor 1 and the second lifting pipe reactor 2.

The first riser reactor 1 and the second riser reactor 2 are each independently selected from one or a combination of more than one of an equal diameter riser reactor, an equal linear velocity riser reactor and a variable diameter riser reactor, and the third reactor 3 is selected from one or a combination of more than one of a fixed fluidized bed reactor, a bulk fluidized bed reactor, a bubbling bed reactor, a turbulent bed reactor, a fast bed reactor, a entrained bed reactor and a dense phase fluidized bed reactor.

The bottom of the first riser reactor 1 is provided with a first hydrocarbon feed line 11 and a first pre-lift gas line 12 for providing a first hydrocarbon feed and introducing a lift gas, respectively. The upper part of the first riser reactor 1 is provided with a recycle oil pipeline 13, and the recycle oil pipeline 13 is positioned at a position 0-1/4 of the distance from the top of the first riser reactor and is used for introducing recycle oil slurry generated by the device.

The bottom of the second riser reactor 2 is provided with a second hydrocarbon feed line 21 and a second pre-lift gas line 22 for providing a second hydrocarbon feed and introducing a lift gas, respectively. The second riser reactor 2 further comprises a second hydrocarbon feedstock line 23 for providing a second hydrocarbon feedstock at another location. Wherein the second hydrocarbon feed line 21 is located at the bottom of the second riser reactor for providing a second hydrocarbon feed having a carbon number of 5 and below, and the second hydrocarbon feed line 23 is located at a position 1/3 to 2/3 of the middle of the second riser reactor from the top for providing a second hydrocarbon feed having a carbon number of greater than 5.

The top outlet of the first riser reactor 1 is directly connected to the bottom of the third reactor 3, while the top outlet of the second riser reactor 2 is connected to the middle upper part of the third reactor 3.

In the catalytic conversion apparatus of the present invention, the combined reactor is arranged in series with the stripper 5, specifically, the stripper 5 is located below the third reactor 3 and communicates with the bottom of the third reactor 3, whereby the water vapor supplied into the stripper 5 can be reused upwardly through the third reactor 3 as the water injection vapor of the catalytic conversion reaction occurring therein, which is advantageous in reducing the overall energy consumption of the reaction.

The stripper 5 is provided with a stripping baffle 55 for reducing the falling speed of the spent catalyst and making the distribution of the stripping gas more uniform, so that the residual reaction oil gas on the spent catalyst is fully removed from the spent catalyst.

In the catalytic conversion device of the present invention, the regenerator 7 is communicated with the stripper 5, and the spent catalyst from the stripper 5 enters the coke-burning tank 6 of the regenerator 7 through the spent catalyst conveying pipe 52 for coke-burning regeneration, so that the spent catalyst is converted into regenerated catalyst. Regenerated catalyst in the regenerator 7 is returned to the pre-riser sections of the first riser reactor 1 and the second riser reactor 2 via a first regenerated catalyst line 72 and a second regenerated catalyst line 71, respectively, for recycling.

In the catalytic conversion apparatus of the present invention, the spent catalyst transfer pipe 51 is connected to the stripper 5 and the second riser reactor 2, respectively, preferably to the middle position of the second riser reactor 2, so that a part of the spent catalyst is transferred from the stripper 5 to the second riser reactor 2, and the transfer rate of the spent catalyst can be adjusted by the slide valve on the spent catalyst transfer pipe 51.

In order to adjust the flow rate of the lift gas and thus flexibly control the reaction environment of the reactor, a lift gas distribution ring 25 may be provided in the middle of the second riser reactor 2, which additionally introduces the lift gas into the second riser reactor 2 through a lift gas line 24. The lifting gas distribution ring 25 is preferably disposed below the connection point of the spent catalyst transfer pipe 51 and the second riser reactor 2 so that the spent catalyst transferred by the spent catalyst transfer pipe 51 can be continuously lifted into the third reactor 3. The lifting gas used is well known to those skilled in the art and may be selected from one or more of steam, nitrogen, dry gas, preferably steam.

A lift gas distribution ring 32 may also be provided at the bottom of the third reactor 3, which introduces a lift gas into the third reactor 3 through a lift gas line 31, thereby adjusting the properties of the catalyst bed in the third reactor 3, such as bed height, bed density, bed space velocity, etc., by the flow rate of the lift gas. Fig. 2 is a schematic cross-sectional view of the lift gas distribution ring 32 of fig. 1. As shown in fig. 2, the lift gas distribution ring 32 has an annular configuration including a plurality of lift gas inlet lines 321 disposed outside the ring and a plurality of lift gas distribution ports 322 disposed inside the ring. The lifting gas used is well known to those skilled in the art and may be selected from one or more of steam, nitrogen, dry gas, preferably steam.

In the catalytic conversion device of the present invention, a settler 4 may be further provided above the third reactor 3 and the stripper 5, the settler 4 being in communication with both the third reactor 3 and the stripper 5, and a separation device, preferably a fast separation device, being provided in the settler 4 and being provided in an upper part of the settler 4, such as the

cyclones

41, 42. The oil gas is rapidly separated from the reacted carbon deposition catalyst through the separation device, so that the yield of dry gas can be reduced, and the conversion of propylene and butene after the generation is inhibited.

After the oil mixture generated in the third reactor 3 enters the settler 4, a small amount of catalyst entrained in the oil mixture is separated through the cyclone separators 41-42, stripping steam in the stripper can also directly enter the settler 4 to be separated together with other oil gas through the cyclone separators 41-42, and the separated cracked oil gas 44 enters the gas collection chamber 43 and then is led out of the combined reactor to enter a subsequent product separation system (not shown in the figure). Cracked gas 44 is separated into products such as dry gas, cracked gas, gasoline, light oil, and slurry oil in a product separation system. The cracked gas can be separated and refined to obtain a polymerization grade propylene product and a C4 fraction rich in olefin. The gasoline is firstly cut into light and medium gasoline fraction segments, and the C4 fraction rich in olefin is partially or completely returned to the second riser reactor 2 for reaction, and the light gasoline is partially or completely returned to the second riser reactor 2 for reaction.

The catalytic conversion method and the catalytic conversion device adopt the combined reactor combining the double lifting pipes and the bed reactor, can provide good reaction environment for the primary cracking of heavy hydrocarbon raw materials and the secondary cracking reaction of gasoline fractions, and can effectively control the properties of the catalysts in the second lifting pipe reactor and the third reactor, such as catalyst activity, metal content, temperature and the like by introducing the spent catalyst into the second lifting pipe reactor.

In addition, the catalytic conversion method and the catalytic conversion device can additionally introduce lifting gas into the second riser reactor and/or the third reactor, and the reaction environment of the reactor can be flexibly controlled by adjusting the flow of the lifting gas, so that the secondary cracking of the gasoline fraction is further promoted.

Through the mode, the catalytic conversion method and the catalytic conversion device can realize higher hydrocarbon conversion capability and higher low-carbon olefin yield.

The present invention will be described in detail with reference to examples, but the scope of the present invention is not limited thereto.

Examples

Reagents, instruments and tests

In the embodiment and the comparative example, the gas product is tested by adopting a petrochemical analysis method RIPP 77-90 method, the coke content is measured by adopting a petrochemical analysis method RIPP 107-90 method, the composition of the organic liquid product is measured by adopting an SH/T0558-1993 method, the cut points of the fractions of gasoline and diesel oil are 221 ℃ and 343 ℃ respectively, and the light aromatic hydrocarbon in the gasoline is measured by adopting a petrochemical analysis method RIPP 82-90.

In the examples below, the conversion of the feedstock and the yield of cracked products were calculated according to the following formulas:

the RIPP petrochemical analysis method used in the present invention is selected from the group consisting of "petrochemical analysis method (RIPP test method)", code Yang Cuiding, scientific Press, 1990.

The reagents used hereinafter are all chemically pure reagents unless otherwise specified.

The Y-type molecular sieves are produced by Qilu catalyst factories and have the following industrial marks:

DASY, physical parameters are: siO (SiO) ₂ /Al ₂ O ₃ ＝50，Na ₂ The O content was 0.85 wt%;

the beta molecular sieve is produced by Qilu catalyst factories; the industrial marks are as follows:

the physical parameters of ZBP are as follows: siO (SiO) ₂ /Al ₂ O ₃ ＝50，Na ₂ The O content was 0.55% by weight.

The MFI structure molecular sieves are all produced by Qilu catalyst factories and have the industrial marks:

ZRP-1: wherein SiO is ₂ /Al ₂ O ₃ ＝30，Na ₂ O content of 0.17 wt%, rare earth oxide RE ₂ O ₃ Wherein the lanthanum oxide content is 0.84 wt.%, the cerium oxide content is 0.18 wt.%, and the other rare earth oxides content is 0.38 wt.%.

The raw materials used in the examples and comparative examples are Daqing wax oil, and specific properties are shown in Table 1.

The catalysts used in the examples and comparative examples were homemade catalysts, denoted CAT, and the active components of the catalysts were Y molecular sieve (DASY), beta molecular sieve, and ZRP molecular sieve, the specific properties of which are shown in table 2.

The catalyst CAT is prepared by the following steps:

uniformly mixing a DASY molecular sieve, a beta molecular sieve and a ZRP molecular sieve, adding deionized water for pulping, and uniformly stirring to obtain molecular sieve slurry with the solid content of 20-40 wt%;

mixing clay, binder and deionized water, pulping, and stirring uniformly to obtain carrier slurry with the solid content of 15-25 wt%;

and mixing and pulping the homogenized molecular sieve slurry and the homogenized carrier slurry, and then sequentially performing spray drying, washing, filtering and drying to obtain the catalyst CAT.

The catalyst CAT was aged at 790℃for 14 hours under 100% steam conditions before the test.

TABLE 1 Properties of the feedstock (Daqing wax oil)

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TABLE 2 composition and Properties of catalyst CAT

Project	CAT
		Elemental composition/% (w)
Al ₂ O ₃	52.8
		SiO ₂	41.2
Microreaction/% (W)	65
		Specific surface area/(m) ² /g)	148
Pore volume/(ml/g)	0.512
		Heap ratio/(g/ml)	0.91
Particle size distribution (%)
		0-20μm	0
0-40μm	16.5
		0-80μm	67.3
0-105μm	89.6
		＞105μm	10.4

Example 1

The test was performed on a medium-sized test apparatus as shown in fig. 1. The apparatus comprises two riser reactors and a fluidized bed reactor. The first riser reactor 1 has an inner diameter of 16mm and a length of 3200mm, the second riser reactor 2 has an inner diameter of 16mm and a height of 3800mm, and the third reactor 3 has an inner diameter of 64mm and a height of 500mm.

Introducing Daqing wax oil into the bottom of the first riser reactor 1, contacting with regenerated catalyst CAT from the regenerator 7 and reacting, and introducing the reacted oil mixture into the third reactor 3;

introducing C4 hydrocarbon into the bottom of the second riser reactor 2, contacting and reacting with regenerated catalyst CAT from a regenerator 7, introducing spent catalyst from the middle of the second riser reactor 2, introducing lifting steam, and introducing the reacted oil mixture into a third reactor 3;

the oil mixture from the first riser reactor 1 and the oil mixture from the second riser reactor 2 are reacted in the third reactor 3, the reacted oil mixture is separated by a cyclone separator, the catalyst enters the stripper 5 and then enters the regenerator 7 for regeneration, the regenerated catalyst returns to the riser reactor for recycling, and the oil gas is introduced into the fractionation system for separation. Wherein the mass ratio of the C4 hydrocarbon to the Daqing wax oil is 0.08:1. The reaction conditions and results are shown in Table 3.

Example 2

The procedure of example 1 was followed except that the C4 hydrocarbon fraction obtained by fractionation was not introduced into the second riser reactor 2, but the light gasoline fraction obtained by fractionation (distillation range 40 to 80 ℃ C., olefin content 65% by weight) was introduced into the second riser reactor 2, and the mass ratio of the light gasoline fraction to Daqing wax oil was 0.08:1. And the spent catalyst and the lifting steam are not introduced in the middle of the second riser 2, but the lifting steam is introduced at the bottom of the third reactor. The reaction conditions and results are shown in Table 3.

Example 3

The procedure of example 1 was followed except that, in addition to the introduction of the fractionated C4 fraction into the second riser reactor 2, a fractionated light gasoline fraction (distillation range 40-80 ℃ C., olefin content 65% by weight) was introduced into the second riser reactor 2, the mass ratio of the C4 hydrocarbons, light gasoline fraction to Daqing wax oil being 0.05:0.05:1. And introducing lift steam at the bottom of the third reactor. The reaction conditions and results are shown in Table 3.

Example 4

The procedure of example 3 was followed, except that the spent catalyst was introduced in the middle of the second riser 2, without introducing lift steam. The reaction conditions and results are shown in Table 3.

TABLE 3 reaction conditions and reaction results for examples 1-4

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Comparative example 1

The test was performed on a medium-sized test apparatus as shown in fig. 1. The apparatus comprises two riser reactors and a fluidized bed reactor. The first riser reactor 1 had an inner diameter of 16mm and a length of 3200mm, the second riser reactor 2 had an inner diameter of 16mm and a height of 3800mm, and the third reactor 3 had an inner diameter of 64mm and a height of 500mm.

Daqing wax oil is introduced into the bottom of the first riser reactor 1, contacts with regenerated catalyst CAT from the regenerator 7 and reacts, and the reacted oil mixture is introduced into the third reactor 3;

The C4 hydrocarbon is introduced into the bottom of the second riser reactor 2, contacts and reacts with the regenerated catalyst CAT from the regenerator 7, and the reacted oil mixture is introduced into the third reactor 3;

the oil mixture from the first riser reactor 1 and the oil mixture from the second riser reactor 2 are reacted in the third reactor 3, the reacted oil mixture is separated by a cyclone separator, the catalyst enters the stripper 5 and then enters the regenerator 7 for regeneration, the regenerated catalyst returns to the riser reactor for recycling, and the oil gas is introduced into the fractionation system for separation. Wherein the indicated mass ratio of the C4 hydrocarbon to the Daqing wax oil is 0.08:1. The reaction conditions and the results are shown in Table 4.

Comparative example 2

The procedure of comparative example 1 was followed except that the C4 hydrocarbon fraction obtained by fractionation was not introduced into the second riser reactor 2, but the light gasoline fraction obtained by fractionation (distillation range 40 to 80 ℃ C., olefin content 65% by weight) was introduced into the second riser reactor 2, and the mass ratio of the light gasoline fraction to Daqing wax oil was 0.08:1. The reaction conditions and the results are shown in Table 4.

Comparative example 3

The procedure of comparative example 1 was followed except that, in addition to the introduction of the fractionated C4 fraction into the second riser reactor 2, a fractionated light gasoline fraction (40-80 ℃ in distillation range, 65 wt.% olefin content) was introduced into the second riser reactor 2, the mass ratio of C4 hydrocarbons, light gasoline fraction to Daqing wax oil being 0.05:0.05:1. The reaction conditions and the results are shown in Table 4.

Comparative example 4

The procedure of comparative example 1 was followed except that, in addition to the introduction of the fractionated C4 fraction into the second riser reactor 2, a fractionated light gasoline fraction (40-80 ℃ in distillation range, 65 wt.% olefin content) was introduced into the second riser reactor 2, the mass ratio of C4 hydrocarbons, light gasoline fraction to Daqing wax oil being 0.08:0.08:1. The reaction conditions and the results are shown in Table 4.

TABLE 4 Table 4

As can be seen from tables 3 and 4, the method and apparatus of the present invention can achieve higher hydrocarbon conversion capability and higher low-carbon olefin yield than the comparative examples.

It will be appreciated by persons skilled in the art that the embodiments described herein are merely exemplary and that various other alternatives, modifications and improvements may be made within the scope of the invention. Thus, the present invention is not limited to the above-described embodiments, but only by the claims.

Claims

1. A catalytic conversion process for producing light olefins comprising:

Mixing the first oil mixture and the second oil mixture, then entering a third reactor for reaction, converting the cracking catalyst into a spent catalyst after reaction, entering a stripper,

wherein a portion of the spent catalyst is introduced into the second riser reactor;

the method further comprises the step of introducing lifting gas into the middle part of the second riser reactor, wherein the lifting gas inlet is positioned below the to-be-regenerated catalyst inlet, and the lifting gas is selected from one or more of water vapor, nitrogen and dry gas;

the reaction conditions of the first riser reactor include: the reaction temperature is 480-620 ℃; the ratio of the agent to the oil is 2-25; the reaction time is 1-15 seconds;

the reaction conditions of the second riser reactor include: the reaction temperature is 560-720 ℃, the catalyst-oil ratio is 3-40, and the reaction time is 0.5-10 seconds;

the reaction temperature of the third reactor is 520-700 ℃ and the weight hourly space velocity is 1-30 hours ^-1 The catalyst density is 100-500 kg/m ³ The height of the bed layer is 1/2-4/5 of the height of the bed layer reactor, and the pressure in the reactor is 0.1-0.4 MPa.

2. The catalytic conversion process according to claim 1, wherein the reaction conditions of the first riser reactor comprise: the reaction temperature is 520-600 ℃; the agent-oil ratio is 3-20; the reaction time is 2-10 seconds.

3. The catalytic conversion process according to claim 1, wherein the reaction conditions of the second riser reactor comprise: the reaction temperature is 580-680 ℃, the catalyst-to-oil ratio is 5-30, and the reaction time is 1-5 seconds.

4. The catalytic conversion process according to claim 1, wherein the first hydrocarbon feedstock is selected from one or more of vacuum wax oil, atmospheric residuum, vacuum residuum, coker wax oil, deasphalted oil, furfural refined raffinate oil, coal liquefied oil, oil sand oil, shale oil, distillate oil obtained by F-T synthesis, and animal and vegetable oils.

5. The catalytic conversion process according to claim 1, wherein the second hydrocarbon feedstock is a mixture of C4 to C8 hydrocarbons, the mixture of C4 to C8 hydrocarbons having an olefin content of greater than 50wt%.

6. The catalytic conversion process according to claim 5, wherein the mixture of C4-C8 hydrocarbons has an olefin content of greater than 60wt%.

7. The catalytic conversion process according to claim 1, wherein the cracking catalyst comprises a cracking active component, clay and a binder, wherein the cracking active component comprises a molecular sieve having an MFI structure,

Wherein, the content of the cracking active component is 20 to 70 weight percent based on the dry weight of the cracking catalyst; the clay content is 15-60 wt%; the content of the binder is 20-35 wt%.

8. The catalytic conversion process of claim 7 wherein the cracking active component is present in an amount of from 30 to 50wt%, based on the dry weight of the cracking catalyst; the clay content is 30-50wt%; the content of the binder is 20-30wt%.

9. The catalytic conversion process of claim 7 wherein the cracking active component further comprises a Y molecular sieve and a beta molecular sieve, wherein the Y molecular sieve is present in an amount of 0 to 90wt%, based on the total weight of the cracking active component; the content of the molecular sieve with the MFI structure is 1-50wt%, and the content of the beta molecular sieve is 0-50wt%.

10. The catalytic conversion process of claim 9 wherein the cracking active component further comprises a Y molecular sieve and a beta molecular sieve, wherein the Y molecular sieve is present in an amount of 50 to 80wt% based on the total weight of the cracking active component; the content of the molecular sieve with the MFI structure is 10-40 wt%, and the content of the beta molecular sieve is 10-40 wt%.

11. The catalytic conversion process according to any one of claims 1 to 9, wherein the lift gas is steam.

12. The catalytic conversion process according to any one of claims 1 to 9, further comprising introducing a lift gas to the third reactor, wherein the lift gas is selected from one or more of steam, nitrogen, dry gas.

13. The catalytic conversion process according to claim 12, wherein the lift gas is steam.

14. The catalytic conversion process according to claim 12, wherein a lift gas is introduced to the bottom of the third reactor.

15. The catalytic conversion process according to any one of claims 1 to 9, further comprising introducing the remaining spent catalyst into a regenerator for regeneration to obtain regenerated catalyst and returning the regenerated catalyst to the first riser reactor and the second riser reactor, respectively.

16. A catalytic conversion device for producing light olefins for use in the catalytic conversion process of claim 1, comprising:

17. The catalytic converter of claim 16, wherein a riser gas distribution ring is provided in the middle of the second riser reactor.

18. The catalytic converter of claim 17, wherein the lift gas distribution ring is disposed below a junction of the spent catalyst transfer pipe and the second riser reactor.

19. The catalytic converter device of claim 16, wherein the bottom of the third reactor is provided with a lift gas distribution ring.

20. The catalytic conversion device of claim 16, further comprising a settler in communication with the third reactor.

21. The catalytic converter of claim 20, further comprising a separation device disposed in an upper portion of the settler.

22. The catalytic conversion device of any one of claims 16 to 21, wherein the first riser reactor and the second riser reactor are each independently selected from one or a combination of more than one of an equal diameter riser reactor, an equal linear velocity riser reactor and a variable diameter riser reactor, and the third reactor is selected from one or a combination of more than one of a fixed fluidized bed reactor, a bulk fluidized bed reactor, a bubbling bed reactor, a turbulent bed reactor, a rapid bed reactor, a entrained bed reactor and a dense fluidized bed reactor.