CN110950731B

CN110950731B - Catalytic cracking method

Info

Publication number: CN110950731B
Application number: CN201811132841.2A
Authority: CN
Inventors: 崔守业; 白旭辉; 李泽坤
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: 2018-09-27
Filing date: 2018-09-27
Publication date: 2022-08-09
Anticipated expiration: 2038-09-27
Also published as: CN110950731A

Abstract

The invention relates to a catalytic cracking method, which comprises the following steps: (1) feeding the catalytic cracking raw oil into a first reaction zone of a catalytic cracking reactor to contact with a catalytic cracking catalyst and carrying out a first catalytic cracking reaction; (2) continuously carrying out a second catalytic cracking reaction on the obtained first reaction product and the semi-spent catalyst in a second reaction zone of the catalytic cracking reactor to obtain a second reaction product and a spent catalyst; and (3) feeding the oxygen-containing compound into a second reaction zone of the catalytic cracking reactor to contact with the first reaction product and the semi-spent catalyst and perform the second catalytic cracking reaction together. The method of the invention can improve the heat distribution of the catalytic cracking reactor and improve the distribution of reaction products.

Description

Catalytic cracking method

Technical Field

The present invention relates to a catalytic cracking process.

Background

Low carbon olefin (C) ₂ -C ₄ Olefins) have been the dominant primary organic chemicals in the modern petroleum and chemical industries, especially ethylene and propylene. The methods for producing light olefins can be broadly divided into two major categories, namely the traditional petroleum route and the emerging non-petroleum route. The traditional method for preparing low-carbon olefin by petroleum route mainly comprises steam cracking and catalytic cracking process production. Since the 10 s in the 20 th century, various countries in the world have been dedicated to developing routes for preparing low-carbon olefins from non-petroleum resources, and some progress has been made.

Chinese patent CN92109905 discloses a catalyst and a reaction process for converting methanol into light olefin, wherein a ZSM-5 type zeolite catalyst containing phosphorus, rare earth elements and a pore structure regulator and a reaction process of a multistage adiabatic fixed bed cracking reactor adopting a dehydration reactor and 2-n reaction-regeneration switching operations are adopted, the non-cyclic operation is carried out at high temperature (more than 400 ℃), and the catalyst has high activity, high selectivity, high water resistance, high thermal stability and long reaction life. On a plant scale of 0.7 to 1 ton of methanol per day, the methanol conversion is 100%, C ₂ -C ₄ The olefin selectivity can be greater than 85%, the on-stream time can be greater than 600 hours, and the once-through operating period can be greater than 24 hours. However, because the heat transfer of the fixed bed reactor is slow, the reaction of preparing light olefins from methanol is a strong exothermic reaction, and hot spots are easy to occur, so that the device is damaged.

Chinese patent CN200810043239.1 discloses a method for producing low-carbon olefin from methanol, which mainly solves the problem of low selectivity of target products in the process of preparing low-carbon olefin from methanol. The method comprises the following steps: (a) heating an oxygen-containing compound raw material containing methanol at a reaction temperature of 300-600 ℃ at a raw material weight hourly space velocity of the oxygen-containing compound of 1-50 hours ^-1 Under the condition that the reaction pressure (gauge pressure) is 0.05-10 MPa, the raw materials are contacted with a silicoaluminophosphate molecular sieve catalyst in a reactor; (b) separating the catalyst from the reaction products; (c) exchanging heat between the reaction product and the raw material containing the methanol, and heating the raw material containing the methanol to 100-350 ℃ under the condition of ensuring effective recovery of heat of the reaction product; (d) repeating the technical schemes (a) to (c).

WO2006049864 discloses a process and corresponding apparatus for the production of light olefins from oxygenates, wherein the process comprises feeding an oxygenate feed stream through a feed stream distributor (3) into an OTO reactor; contacting the oxygenate with a catalyst to produce a mixture comprising light olefins, unreacted oxygenate, and other by-products; separating unreacted oxygenate and diolefins from said light olefins and said by-products; and returning unreacted oxygenate and diolefins to the OTO reactor. Unreacted oxygenate and diolefins are fed to the reactor through at least one feed nozzle at a point separate from the oxygenate feed stream. The process is believed to be able to carry out the oxygenate conversion over a wide range of pressures (10.1kPa to 10.1MPa), but not all pressures give good results, preferably from 101.3kPa to 1013.3 kPa.

After the dehydration reaction of the oxygen-containing compound to prepare the low-carbon olefin, a certain amount of water is also by-produced in addition to the hydrocarbon product, in the case of methanol, about 44% of the hydrocarbon product can be obtained after the reaction, and in the case of ethanol, 46% of the hydrocarbon product can be obtained after the reaction. It is known that the reaction of an oxygenate feedstock to produce lower olefins is a reaction in which the number of molecules increases, and that low reaction pressures favor chemical equilibrium toward the production of lower olefins. In view of this, in the prior art, lower reaction pressure is generally used to obtain the desired yield of lower olefins. The direct consequence of this low reaction pressure (typically 0.1-0.3MPa) is that if it is desired to increase the oxygenate feedstock throughput for the purpose of increasing the production of lower olefins, the prior art has had to increase the size or number of reactors for this purpose in order to maintain the yield of lower olefins at an acceptable level. Obviously, this would correspondingly increase the investment and maintenance costs of the equipment.

Because the reaction process of preparing the low-carbon olefin from the oxygen-containing compound is an exothermic reaction, if pure methanol is used for feeding, the total reaction heat of the methanol-to-olefin is generally 20-35 KJ/mol, the adiabatic temperature rise is above 200 ℃ at a small water-alcohol ratio, and if a side reaction in the MTO reaction process is considered, the adiabatic temperature rise is larger. Such high temperature rise not only affects the MTO reaction result and accelerates the carbon deposition rate of the catalyst, but also needs to consider the hydrothermal stability of the catalyst. Therefore, to reduce the temperature rise in the reactor, measures such as designing a heat removal system, reducing the initial composition of reactants, and reducing the temperature of the feed are generally required.

Chinese patent CN201110253681.9 discloses a method for reducing energy consumption of an apparatus for producing olefin from oxygen-containing compound. Preheating a liquid oxygen-containing compound raw material to a certain temperature by a raw material preheater, dividing the raw material into two strands, and heating and gasifying one strand by a raw material vaporizer to obtain a gas-phase raw material; the other is atomized into atomized liquid phase raw material; mixing a gas-phase raw material and an atomized liquid-phase raw material in front of a raw material/reaction gas heat exchanger, feeding the mixture into the raw material/reaction gas heat exchanger in a mist flow mode, fully exchanging heat with oil gas generated by high-temperature reaction from a reactor to recover high-temperature-level heat of the oil gas generated by the high-temperature reaction, completely gasifying the raw material after recovering the high-temperature-level heat to form high-temperature gasification raw material gas, and feeding the high-temperature gasification raw material gas into the reactor for reaction; oil gas generated by high-temperature reaction from the reactor is sent to a rear quenching water washing system after heat exchange of the raw material/reaction gas heat exchanger. The method can effectively improve the heat exchange effect of the heat exchanger, reduce the volume of the heat exchanger and reduce the energy consumption of the device, but the method does not solve the problem of cooling the reaction catalyst.

At present, the reactors adopted by the technology for preparing low-carbon olefin by using oxygen-containing compounds mainly comprise a fixed fluidized bed reactor, a dense-phase fluidized bed reactor and the like. The fixed bed reactor has poor bed heat transfer effect, and for strong exothermic reaction, if the heat released in the reaction process can not be transferred out in time, hot spots, temperature runaway and other problems are easy to occur, and the replacement and regeneration of the catalyst are relatively troublesome. The diameter and height of the fixed bed reactor are not strictly limited, but the height/diameter ratio adopted by the reactor design is 2.5-12 in consideration of the factors of fluid distribution, manufacturing cost, safety and the like. The fixed bed reactor is also adopted in the early catalytic cracking reaction, carbon deposition can occur on the catalyst in the catalytic cracking reaction, and the regeneration reaction is needed to recover the activity of the catalyst, so that the fixed bed reactor is needed to be discontinuously used for reaction and regeneration, and at least more than three reactors are needed to ensure the continuity of continuous feeding and other operations. The dense-phase fluidized bed reactor belongs to the category of bubbling bed and turbulent bed, the reaction residence time is generally longer, the residence time of the generated olefin in the reactor is longer, the hydrogen transfer reaction is increased, and the dense-phase fluidized bed reactor is also very unfavorable for the high yield of low-carbon olefin. Due to the low linear velocity of the dense phase fluidized bed reactor, back mixing tends to occur, affecting product distribution and quality, and on the other hand the diameter of the reactor is relatively large.

Since the birth of the 40 th generation of the 20 th century, catalytic cracking has been the most important process for lightening heavy oil. The reason is that the raw material source is wide, wax oil can be adopted, deasphalted oil of atmospheric residue oil and vacuum residue oil can also be adopted, or vacuum residue oil can be partially mixed; secondly, the product scheme is flexible, and can be fuel type or fuel chemical type, such as productive gasoline, productive diesel oil, productive propylene and the like; and thirdly, the product properties can be correspondingly adjusted through adjustment of a catalyst formula and change of process parameters, such as improvement of the octane number of the gasoline, reduction of the olefin content of the gasoline and the like.

Chinese patent CN99105904.2 discloses a catalytic conversion method for preparing isobutane and gasoline rich in isoparaffin, which comprises the steps of enabling preheated raw oil to enter a reactor comprising two reaction zones, and contacting the preheated raw oil with a hot cracking catalyst, wherein the temperature of the first reaction zone is 530-620 ℃, and the reaction time is 0.5-2.0 seconds; the temperature of the second reaction zone is 460-530 ℃, the reaction time is 2-30 seconds, the reaction product is separated, and the catalyst to be generated enters a regenerator through steam stripping and is recycled after being burnt. The liquefied gas prepared by the method provided by the patent contains 20-40 wt% of isobutane, 30-45 wt% of isoparaffin in gasoline family components, and the content of olefin is reduced to be below 30 wt%, wherein the research octane number is 90-93, and the motor octane number is 80-84.

Chinese patent CN99105905.0 discloses a catalytic conversion method for preparing propylene, isobutane and gasoline rich in isoparaffin, which comprises the steps of enabling preheated raw oil to enter a reactor comprising two reaction zones, and contacting the preheated raw oil with a hot cracking catalyst, wherein the temperature of the first reaction zone is 550-650 ℃, and the reaction time is 0.5-2.5 seconds; the temperature of the second reaction zone is 480-550 ℃, the reaction time is 2-30 seconds, the reaction product is separated, and the catalyst to be generated enters a regenerator through steam stripping and is burnt for recycling. The yield of the liquefied gas prepared by the method provided by the patent can reach 25-40 wt%, wherein the content of propylene is about 30 wt%, the content of isobutane is 20-40 wt%, the yield of gasoline can reach 35-50 wt%, and the content of isoparaffin in the gasoline is 30-45 wt%.

Chinese patent CN99105903.4 discloses a riser reactor for fluidized catalytic conversion, which comprises a pre-lifting section, a first reaction zone, a second reaction zone with enlarged diameter, and an outlet zone with reduced diameter, which are coaxial with each other, from bottom to top along the vertical direction, and a horizontal pipe is arranged at the end of the outlet zone. The reactor can control the technological conditions of the first reaction zone and the second reaction zone to be different, and can also carry out staged cracking on raw oil with different properties to obtain a required target product.

The heat required by the current catalytic cracking reactor is mainly provided by the heat regenerated catalyst circulated back, and the purpose can be achieved only by increasing the heat at the bottom of a riser in order to increase the reaction temperature, but the defects are that the contact temperature of the catalyst and the oil is too high, or the catalyst-oil ratio is too large, so that the yield of dry gas and coke is obviously increased, and the product distribution is poor. In addition, because the riser reactor dissipates heat, the raw material is vaporized, the hydrocarbon reacts, the temperature of the riser reactor is obviously reduced along with the increase of the height of the riser, particularly when the variable-diameter riser is adopted, the temperature of the secondary reaction is reduced by 10-15 ℃ compared with that of the primary reaction, and the cracking reaction of the secondary reaction is unfavorable.

Disclosure of Invention

The invention aims to provide a catalytic cracking method, which can improve the heat distribution of a catalytic cracking reactor and improve the distribution of reaction products.

The invention provides a catalytic cracking method, which comprises the following steps:

(1) feeding the catalytic cracking raw oil into a first reaction zone of a catalytic cracking reactor to contact with a catalytic cracking catalyst and perform a first catalytic cracking reaction to obtain a first reaction product and a semi-finished catalyst; the catalytic cracking reactor sequentially comprises a first reaction zone and a second reaction zone which are communicated by fluid according to the flow direction of reaction materials;

(2) continuously carrying out a second catalytic cracking reaction on the obtained first reaction product and the semi-spent catalyst in a second reaction zone of the catalytic cracking reactor to obtain a second reaction product and a spent catalyst, and sending the second reaction product and the spent catalyst out of the catalytic cracking reactor through an outlet of the second reaction zone;

(3) feeding the obtained spent catalyst into a regenerator for regeneration, and returning the obtained regenerated catalyst serving as the catalytic cracking catalyst to the catalytic cracking reactor;

the method further comprises the following steps: in the step (2), an oxygen-containing compound is fed into a second reaction zone of the catalytic cracking reactor through an oxygen-containing compound inlet to contact with the first reaction product and the semi-spent catalyst and carry out a second catalytic cracking reaction together; the temperature at the outlet of the second reaction zone is controlled to be 0.3 to 100 ℃ higher than the temperature at the inlet of the oxygenate (the temperature of the reaction zone).

In some embodiments of the invention, the first reaction zone and the second reaction zone are divided by the location of oxygenate injection, with the reaction zone upstream of the location of oxygenate injection being the first reaction zone and the reaction zone downstream of the location of oxygenate injection being the second reaction zone, in terms of reactant feed flow direction.

In some embodiments of the invention, the invention optimizes product selectivity by feeding an oxygenate to the reactor to improve reaction conditions in the latter half of the reactor. Catalytic cracking reactions are well known to those skilled in the art, and reference may be made specifically to "catalytic cracking chemistry and technology", written to school, 2013 edition, chapter v. The conditions of the first catalytic cracking reaction may include: the temperature (the temperature of the oxygen-containing compound is fed into the reactor) is 450-620 ℃, preferably 500-600 ℃, the reaction time is 0.5-2 seconds, preferably 0.8-1.5 seconds, the weight ratio of the catalytic cracking catalyst to the catalytic cracking raw oil is 3-15, preferably 3-12, the reaction pressure is 130-450 kPa, and the weight ratio of the steam to the catalytic cracking raw oil is 0.03-0.3; the conditions of the second catalytic cracking reaction may include: the temperature (temperature at the outlet of the second reaction zone) is 460-660 ℃, preferably 505-620 ℃ and the reaction time is 2-30 seconds, preferably 3-15 seconds.

In some embodiments of the invention, the catalytic cracking reaction of the oxygenate produces water, which contributes to a reduction in the partial pressure of the reacting hydrocarbons, the oxygenate undergoing the second catalytic cracking reaction produces water in an amount W1 (based on the total conversion of oxygen in the oxygenate to water), and the weight ratio of W1 to steam injected in the first catalytic cracking reaction may be in the range of from 10 to 200 wt%, preferably from 40 to 150 wt%.

In some embodiments of the present invention, under the same catalyst-to-oil ratio condition, the method of the present invention can reduce the contact temperature of the regenerated catalyst and the raw material by more than 10 ℃, reduce thermal cracking reaction, reduce temperature difference in the second catalytic cracking reaction process, supplement heat required for cracking for the riser reactor, especially for the second reaction zone of the variable diameter riser reactor, reduce the partial pressure of oil gas and hydrocarbon, enhance selective cracking reaction, hydrogen transfer reaction and isomerization reaction, and improve the yield of low carbon olefins such as ethylene and propylene. It has been found through experiments that a better reaction result can be obtained by controlling the temperature at the outlet of the second reaction zone to be 0.3 to 100 ℃, preferably 0.3 to 50 ℃, more preferably 0.5 to 20 ℃, and even more preferably 0.8 to 5 ℃ higher than the temperature at the inlet of the oxygen-containing compound.

In some embodiments of the present invention, it is preferable to control the temperature before the oxygenate is fed to the catalytic cracking reactor so as to control the temperature of the second catalytic cracking reaction, for example, the temperature of the oxygenate fed to the catalytic cracking reactor may be 20 to 600 ℃, preferably 200 to 500 ℃, and more preferably 360 to 450 ℃.

In some embodiments of the present invention, the catalytic cracking feed oil is well known to those skilled in the art, and may be selected from one or more of vacuum wax oil, atmospheric wax oil, coker wax oil, deasphalted oil, furfural refined raffinate oil, coal liquefied oil, oil sand oil, shale oil and animal and vegetable fats and oils, for example, and the carbon content of the semi-spent catalyst may be 0.2 to 3 wt%, preferably 0.2 to 2 wt%, based on the weight of the semi-spent catalyst.

In some embodiments of the present invention, the oxygen-containing compound is well known to those skilled in the art, and refers to an organic compound containing oxygen atoms, which may also be referred to as an organic oxygen-containing compound, and may be one or more selected from alcohols, ethers, and lipids, preferably one or more selected from methanol, ethanol, methyl ether, ethyl ether, methyl ethyl ether, dimethyl carbonate, and methyl formate, and more preferably methanol.

In some embodiments of the present invention, the conditions of the second catalytic cracking reaction may be controlled by adjusting the feed amounts of the oxygenate and the catalytic cracking feedstock, and the ratio of the feed amount of the oxygenate to the feed amount of the catalytic cracking feedstock may be 0.02 to 50, preferably 0.05 to 2.

In some embodiments of the present invention, the catalytic cracking catalyst is well known to those skilled in the art, for example, the catalytic cracking catalyst may include a molecular sieve, which may be one or more selected from the group consisting of a Y-type molecular sieve, a ZSM-5 type molecular sieve, a SAPO type molecular sieve, a beta type molecular sieve, and an SRM type silicoaluminophosphate molecular sieve, a binder, and a matrix, and other conventional catalysts may be used by those skilled in the art.

In some embodiments of the invention, the method may further comprise: at least a portion of the catalytic cracking feed oil is fed into a second reaction zone for the second catalytic cracking reaction, thereby further improving heat distribution.

In some embodiments of the present invention, catalytic cracking reactors are well known to those skilled in the art, for example, the catalytic cracking reactor can be a riser reactor and/or a fluidized bed reactor, and the riser reactor can be one or more selected from the group consisting of a constant diameter riser, a variable diameter riser, and a constant linear velocity riser.

In some embodiments of the present invention, the inventors have assiduously studied and experimented that injecting an oxygen-containing compound at the rear of a riser reactor can increase the heat of the riser reactor; when the reducing riser reactor is adopted, the oxygen-containing compound can obtain a good cracking environment, the heat loss caused by the cracking and heat dissipation of the conventional hydrocarbons in the environment is matched with the heat provided by the oxygen-containing compound, and the water generated by the reaction of the oxygen-containing compound can reduce the hydrocarbon partial pressure of the second catalytic cracking reaction and improve the product distribution.

In a preferred embodiment, the catalytic cracking reactor may be a riser reactor, the first reaction zone is located below the second reaction zone, and the inner diameter of the second reaction zone is greater than the inner diameter of the first reaction zone.

In some embodiments of the invention, the method may further comprise: and (3) taking out at least part of the semi-regenerated catalyst containing carbon from the regenerator in the step (3), and sending the semi-regenerated catalyst into a catalytic cracking reactor to carry out the first catalytic cracking reaction and/or the second catalytic cracking reaction, so that the semi-regenerated catalyst can be used for supplementing reaction heat.

In some embodiments of the present invention, the process provided herein can be carried out in a constant diameter riser, a constant linear velocity riser, or a fluidized bed reactor, wherein the constant diameter riser is the same as a refinery conventional catalytic cracking reactor and the linear velocity of the fluid in the constant linear velocity riser is substantially the same. The constant-diameter riser and the constant-linear-velocity riser reactor sequentially comprise a pre-lifting section, a first reaction zone and a second reaction zone from bottom to top, the fluidized bed reactor sequentially comprises the first reaction zone and the second reaction zone from bottom to top, and the height ratio of the first reaction zone to the second reaction zone is 10-40: 90-60. When an equal-diameter riser, an equal-linear-speed riser or a fluidized bed reactor is used, one or more cold shock medium inlets are arranged at the bottom of the second reaction zone, and/or a heat collector is arranged in the second reaction zone, wherein the height of the heat collector accounts for 50% -90% of the height of the second reaction zone. The temperature and reaction time of each reaction zone are controlled separately. The chilling medium is one or a mixture of more than one of chilling agent, cooled regenerated catalyst and cooled semi-regenerated catalyst in any proportion. Wherein the chilling agent is one or a mixture of more than one of liquefied gas, crude gasoline, stable gasoline, diesel oil, heavy diesel oil or water in any proportion; the cooled regenerated catalyst and the cooled semi-regenerated catalyst are obtained by cooling the spent catalyst after two-stage regeneration and one-stage regeneration respectively, the carbon content of the regenerated catalyst is less than 0.1 wt%, preferably less than 0.05 wt%, and the carbon content of the semi-regenerated catalyst is 0.1-0.9 wt%, preferably 0.15-0.7 wt%.

In some embodiments of the present invention, the process provided by the present invention can also be carried out in a composite reactor consisting of a constant diameter riser and a fluidized bed, wherein the lower constant diameter riser is a first reaction zone and the upper fluidized bed is a second reaction zone, and the temperature and the reaction time of each reaction zone are respectively controlled. One or more chilling medium inlets are arranged at the bottom of the fluidized bed, and/or a heat collector is arranged in the second reaction zone, and the height of the heat collector accounts for 50% -90% of the height of the second reaction zone. The temperature and reaction time of each reaction zone are controlled separately. The chilling medium is one or more than one mixture in any proportion selected from chilling agent, cooled regenerated catalyst and cooled semi-regenerated catalyst. Wherein the chilling agent is one or a mixture of more than one of liquefied gas, crude gasoline, stable gasoline, diesel oil, heavy diesel oil or water in any proportion; the cooled regenerated catalyst and the cooled semi-regenerated catalyst are obtained by cooling the spent catalyst after two-stage regeneration and one-stage regeneration respectively, the carbon content of the regenerated catalyst is less than 0.1 wt%, preferably less than 0.05 wt%, and the carbon content of the semi-regenerated catalyst is 0.1-0.9 wt%, preferably 0.15-0.7 wt%.

In some embodiments of the present invention, the process provided herein can also be carried out in a variable diameter riser reactor (see ZL 99105903.4).

In one embodiment, preheated catalytic cracking raw oil is contacted with a heat regeneration catalyst at the bottom of a conventional equal-diameter riser reactor to perform a cracking reaction in a first reaction zone, generated oil gas and a carbon-containing catalyst ascend to be contacted with an introduced semi-regeneration catalyst, then a cracking reaction, an isomerization reaction and a hydrogen transfer reaction are performed, then an oxygen-containing compound is injected at the bottom of a second reaction zone, the preheated catalytic cracking raw oil is contacted with a carbon-containing semi-spent catalyst and a first reaction product, the oxygen-containing compound is dehydrated to generate low-carbon olefin, and the second reaction product and the spent catalyst enter a settler after the reaction; separating reaction products, feeding the spent catalyst into a two-section regenerator for burning after steam stripping, feeding the semi-regenerated catalyst from the first section regenerator into the middle lower part of the reactor after cooling, and directly returning the regenerated catalyst from the second section regenerator to the bottom of the reactor without cooling.

In another embodiment, for a catalytic cracking device with a reducing riser reactor, preheated catalytic cracking raw oil enters from the lower part of a first reaction zone of the reactor to contact with a heat regeneration catalyst to carry out cracking reaction, a first reaction product and a semi-spent catalyst generated after the reaction move upwards to the lower part of a second reaction zone of the reactor, an injected oxygen-containing compound is dehydrated to generate low-carbon olefin, heat and water vapor are released, the first reaction product and the semi-spent catalyst contact in the second reaction zone to carry out cracking reaction, hydrogen transfer reaction and isomerization reaction, and an effluent after the reaction enters a settler; separating reaction product, stripping catalyst to be regenerated, and regenerating.

Compared with the prior art, the invention has the advantages that:

1. under the condition of the same catalyst-to-oil ratio, the contact temperature of the bottom catalyst (the temperature of the hot catalyst when the hot catalyst is in contact with the raw material) can be reduced by more than 10 ℃, and the thermal cracking reaction is reduced.

2. Under the condition of the same catalyst-oil contact temperature, the catalyst-oil ratio can be below 15, and the generation of coke by the catalyst-oil ratio is reduced.

3. The reaction temperature at the rear part of the oxygen-containing compound injection opening is increased, the heat required by the cracking reaction is supplemented for the second reaction area of the riser reactor, particularly the reducing riser reactor, the oil-gas hydrocarbon partial pressure is reduced, and the selective cracking reaction, the hydrogen transfer reaction and the isomerization reaction are enhanced.

4. The ethylene content in the dry gas produced by the method provided by the invention is increased by more than 0.1 percent, and the propylene content in the liquefied gas is increased by more than 1.0 percent.

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 is a schematic diagram of the structure of one embodiment of a reactor used in the process of the present invention.

Description of the reference numerals

I first reaction zone II second reaction zone

1 line 2 riser reactor 3 lines

4 line 5 line 6 line

7 pipeline 8 settler 9 gas collection chamber

10 stripping section 11 pipeline 12 inclined tube

13 regenerator 14 line 15 line

16 inclined tube

Detailed Description

The following detailed description of embodiments of the invention refers to the accompanying drawings. 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 FIG. 1, a reducing riser reactor is used, and the shape and size of the equipment and lines are not limited by the attached drawings but determined according to specific conditions.

The pre-lifting medium enters a first reaction zone I from the bottom of a riser reactor 2 through a pipeline 1, the regenerated catalyst from an inclined pipe 16 moves upwards in an accelerated manner along the riser reactor under the lifting action of the pre-lifting medium, the reaction raw material is injected into the bottom of the first reaction zone I of the riser reactor 2 through a pipeline 3 and atomized steam from a pipeline 4, the reaction raw material is subjected to a first catalytic cracking reaction on the hot regenerated catalyst and moves upwards in an accelerated manner, and the semi-spent catalyst and a first reaction product are obtained. An oxygen-containing compound is injected into the bottom of a second reaction zone II of the riser reactor 2 together with atomized steam from a pipeline 6 through a pipeline 5, the oxygen-containing compound and a semi-spent catalyst containing certain carbon are dehydrated to generate a second catalytic cracking reaction of olefin, a first to-be-produced substance from a first reaction zone I continues the second catalytic cracking reaction in the second reaction zone II, a generated second reaction product and an inactivated spent catalyst enter a settler 8 through a pipeline 7, the spent catalyst and the second reaction product are separated through a cyclone separator, the second reaction product enters a gas collection chamber 9, and catalyst fine powder returns to the settler through a cyclone separator dipleg. Spent catalyst in the settler flows to the stripping section 10 where it is contacted with steam from line 11. The reaction product oil gas extracted from the spent catalyst enters a gas collection chamber 9 after passing through a cyclone separator. The stripped spent catalyst enters a regenerator 13 through an inclined pipe 12, the main air enters the regenerator through a pipeline 14 to burn off coke on the spent catalyst, the inactivated spent catalyst is regenerated, and the flue gas enters a smoke machine through a pipeline 15. The regenerated catalyst enters the riser reactor through the inclined tube 16. The reaction product oil gas in the gas collection chamber 9 enters a subsequent separation system through a large oil gas pipeline 17.

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

The catalysts used in the examples and comparative examples were prepared as follows:

the molecular sieve slurry is prepared by the steps of adding 22 kg of ZSM-5 molecular sieve (ZSP-2, iron oxide content 2.2 wt%, phosphorus pentoxide content 2.5 wt%, product of Chinese petrochemical catalyst Qilu division) in dry basis (weight basis of molecular sieve, the same below), adding 40 kg of deionized water for pulping, adding 8 kg of USY molecular sieve (DASY2.0, rare earth oxide content 1.8 wt%, unit cell constant 2.445 nm, crystallinity 68%, product of Chinese petrochemical catalyst Qilu division) and then adding 10 kg of SRM molecular sieve (SRM-8, product of Chinese petrochemical catalyst Qilu division) and uniformly pulping. Mixing the aluminum sol (commercially available, Al) ₂ O ₃ Content of 25 wt%) 80 kg of slurry, adding kaolin (product of China Kaolin Corp.) of 40 kg on a dry basis, uniformly pulping, adding the above molecular sieve slurry, uniformly pulping, spray-drying, and collecting microsphere particles in the range of 20-180 micrometers. The catalyst microspheres were calcined at 500 ℃ for 1 hour and then treated with 1 wt% NH ₄ And (3) carrying out exchange treatment on the Cl aqueous solution at 60 ℃ for 0.5 hour, filtering, washing with water, and drying at 120 ℃ for 16 hours to obtain the catalyst A adopted in the embodiment and the comparative example of the invention.

Catalyst a consisted of: 22 wt% of ZSP-2 zeolite, 8 wt% of DASY2.0 molecular sieve, 10 wt% of SRM-8 molecular sieve, 20 wt% of alumina sol (calculated by alumina) and 40 wt% of kaolin.

Catalyst A was aged in a fluidized bed aging apparatus at 800 ℃ under 100% steam for 17 hours prior to the catalytic reaction.

The methanol used in the examples and the comparative examples is pure grade methanol analyzed by Beijing chemical plant, and the methanol content is more than or equal to 99.5 weight percent.

Example 1

Example 1 in a specific embodiment the foregoing procedure is carried out in a medium-sized apparatus as shown in figure 1. The properties of the starting material A used are shown in Table 1, and the catalyst used is catalyst A. The preheating temperature of the oxygen-containing compound was 360 ℃ and the carbon content of the spent catalyst was 0.95% by weight. The reaction conditions and the product distribution are shown in Table 2.

Comparative example 1

Comparative example 1 was carried out in the medium-sized apparatus shown in fig. 1 according to the aforementioned procedure, except that: the raw material A and the methanol respectively enter the riser reactor from the bottoms of the riser reactors of the two sets of medium-sized devices for reaction, and product distribution is respectively obtained. The product distributions of feed A and methanol were multiplied by the respective feed ratios and then summed to obtain the product distribution of comparative example 1. The reaction conditions and the product distribution are shown in Table 2.

The method of the invention can reduce the mixing temperature of the oil in the first reaction area, and can further reduce the coke yield and improve the yields of ethylene, liquefied gas, propylene and butylene as can be seen from the table 2.

The preferred embodiments of the present invention have been described in detail with reference to the accompanying drawings, however, the present invention is not limited to the specific details of the above embodiments, and various simple modifications can 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

Temperature difference: the difference between the temperature at the outlet of the second reaction zone and the bottom of the second reaction zone (i.e., at the oxygenate inlet).

Claims

1. A method of catalytic cracking, the method comprising:

characterized in that the method further comprises: in the step (2), an oxygen-containing compound is fed into a second reaction zone of the catalytic cracking reactor through an oxygen-containing compound inlet to contact with the first reaction product and the semi-spent catalyst and carry out a second catalytic cracking reaction together; the temperature of the oxygen-containing compound sent into the catalytic cracking reactor is 200-450 ℃; controlling the temperature at the outlet of the second reaction zone to be 0.3-100 ℃ higher than the temperature at the inlet of the oxygen-containing compound;

the oxygen-containing compound is one or more selected from alcohols, ethers and lipids.

2. The method of claim 1, wherein the conditions of the first catalytic cracking reaction comprise: the temperature is 450-;

the conditions of the second catalytic cracking reaction include: the temperature is 460-650 ℃, and the reaction time is 2-30 seconds.

3. The process of claim 1, wherein the second catalytic cracking reaction of the oxygenate produces water in an amount W1 and the weight ratio of W1 to the steam injected in the first catalytic cracking reaction is in the range of 10-200 wt%.

4. The process as claimed in claim 3, wherein the weight ratio of W1 to steam injected for the first catalytic cracking reaction is 40-150 wt%.

5. The process of claim 1, wherein the temperature at the outlet of the second reaction zone is 0.3 to 50 ℃ higher than the temperature at the inlet of the oxygenate.

6. A process according to claim 5, wherein the temperature at the outlet of the second reaction zone is between 0.5 and 20 ℃ higher than the temperature at the inlet of the oxygenate.

7. The process as claimed in claim 1, wherein the temperature of the oxygenate fed to the catalytic cracking reactor is 360-450 ℃.

8. The method according to claim 1, wherein the catalytic cracking raw oil is one or more selected from vacuum wax oil, atmospheric wax oil, coker wax oil, deasphalted oil, furfural refined raffinate oil, coal liquefaction oil, oil sand oil, shale oil and animal and vegetable oil.

9. The process of claim 1, wherein the semi-spent catalyst comprises from 0.1 to 3 weight percent char, based on the weight of the semi-spent catalyst.

10. The process of claim 9, wherein the semi-spent catalyst comprises from 0.2 to 2 weight percent char, based on the weight of the semi-spent catalyst.

11. The method according to claim 1, wherein the oxygen-containing compound is one or more selected from methanol, ethanol, methyl ether, ethyl ether, methyl ethyl ether, dimethyl carbonate and methyl formate.

12. The process according to claim 1, wherein the ratio of the feed amount of the oxygenate to the feed amount of the catalytic cracking raw oil is from 0.02 to 50.

13. The method according to claim 12, wherein the ratio of the feed amount of the oxygenate to the feed amount of the catalytic cracking raw oil is 0.05 to 2.

14. The method of claim 1, wherein the catalytic cracking catalyst comprises a molecular sieve, a binder, and a matrix, the molecular sieve being one or more selected from the group consisting of Y-type molecular sieves, ZSM-5 type molecular sieves, SAPO type molecular sieves, beta type molecular sieves, and SRM type silicoaluminophosphate molecular sieves.

15. The process of claim 1, wherein the catalytic cracking reactor is a riser reactor and/or a fluidized bed reactor, and the riser reactor is one or more selected from the group consisting of a constant diameter riser, a variable diameter riser, and a constant linear velocity riser.

16. The process of claim 1, wherein the catalytic cracking reactor is a riser reactor, the first reaction zone is located below the second reaction zone, and the second reaction zone has an inner diameter greater than the inner diameter of the first reaction zone.

17. The method of claim 1, further comprising: at least part of the catalytic cracking raw oil is sent into a second reaction zone to carry out the second catalytic cracking reaction.

18. The method of claim 1, further comprising: and (3) taking out at least part of the semi-regenerated catalyst containing carbon from the regenerator of the step (3), and sending the semi-regenerated catalyst into a catalytic cracking reactor to carry out the first catalytic cracking reaction and/or the second catalytic cracking reaction.