CN102453539B - Catalytic conversion method used for raising gasoline octane number - Google Patents

Abstract

A catalytic conversion method for increasing gasoline octane barrels. High-quality feedstock oil and a thermally regenerated catalyst with low activity are contacted in the lower part of the reactor to undergo a cracking reaction, and the generated oil gas and carbon-containing catalyst ascend under a certain reaction environment. Selective hydrogen transfer reaction and isomerization reaction, separation of reaction products, recycled catalyst after stripping and regeneration; cutting the diesel fraction in the reaction product into light diesel fraction and heavy diesel fraction, heavy diesel fraction alone or /and the back-refined oil fraction and oil slurry fraction are subjected to hydrotreating to obtain a hydrogenated heavy diesel oil fraction, and the hydrogenated heavy diesel oil fraction or/and light diesel oil fraction are returned to this reactor or/and other reactors for further reaction. The method increases the gasoline yield by more than 5% by weight, the octane number by one unit, the isobutene content in the liquefied gas by more than 30% by weight, and the olefin content in gasoline to more than 25% by weight.

Description

A kind of catalytic conversion method for improving gasoline octane number barrel

technical field

The invention belongs to the catalytic conversion method of petroleum hydrocarbons in the absence of hydrogen, more specifically, it belongs to the catalytic conversion method of increasing the gasoline octane number barrel.

Background technique

The conventional catalytic cracking process is mainly used to produce gasoline, and the gasoline yield is as high as 50% by weight. In the early 1980s, unleaded gasoline forced the development of catalytic cracking technology to produce high-octane gasoline. For this reason, the process conditions and catalyst types of catalytic cracking have undergone great changes. In terms of technology, it is mainly to increase the reaction temperature, shorten the reaction time, increase the severity of the reaction, inhibit the hydrogen transfer reaction and overcracking reaction, and improve the contact efficiency between oil gas and catalyst at the bottom of the riser; in terms of catalyst, the USY type zeolite combined with inertness has been developed. Matrix or active matrix catalysts and catalysts composited with different types of zeolites.

Although catalytic cracking technology has made the above progress, met the requirements of unleaded gasoline and increased the octane number of gasoline, whether it is through changing the process conditions or using a new type of zeolite catalyst to increase the octane number of gasoline, it is all based on Increase the olefin content in gasoline components to increase the octane number of gasoline. Currently, the olefin content in gasoline components is 35-65% by weight, which is far from the requirements for olefin content in newly formulated gasoline. The content of olefins in the liquefied gas composition is higher, about 70% by weight, and butene is several times that of isobutane, so it is difficult to be used as an alkylation raw material.

ZL99105904.2 discloses a catalytic conversion method for producing isobutane and isoparaffin-rich gasoline, which is to enter the preheated raw oil into a reactor including two reaction zones, and contact with the hot cracking catalyst , the temperature of the first reaction zone is 530-620°C, the reaction time is 0.5-2.0 seconds; the temperature of the second reaction zone is 460-530°C, the reaction time is 2-30 seconds, the reaction product is separated, and the catalyst to be produced is steamed Put it into the regenerator and burn it for recycling. The content of isobutane in the liquefied gas produced by the method provided by the invention is 20-40% by weight, the content of isoparaffins in the gasoline group composition is 30-45% by weight, and the content of olefins is reduced to below 30% by weight. The French octane number is 90-93, and the motor octane number is 80-84.

ZL99105905.0 discloses a catalytic conversion method for producing propylene, isobutane and gasoline rich in isoparaffins. The preheated raw oil is put into a reactor including two reaction zones, and thermal cracking Catalyst contact, the temperature of the first reaction zone is 550-650°C, the reaction time is 0.5-2.5 seconds; the temperature of the second reaction zone is 480-550°C, the reaction time is 2-30 seconds, the reaction product is separated, and the catalyst to be produced After being stripped, it enters the regenerator and is burned for recycling. The yield of liquefied gas produced by the method provided by the invention can reach 25-40 wt%, wherein the content of propylene is about 30 wt%, the content of isobutane is 20-40 wt%, and the yield of gasoline can reach 35-50 wt%. % by weight, and isoparaffins in the gasoline composition are 30 to 45% by weight.

ZL99105903.4 discloses a riser reactor for fluidized catalytic conversion. Along the vertical direction from bottom to top, there are mutually coaxial pre-lift section, first reaction zone, second reaction zone with enlarged diameter, A reduced diameter outlet zone with a horizontal tube at the end of the outlet zone. The reactor can not only control the difference in process conditions between the first reaction zone and the second reaction zone, but also make the feedstock oils with different properties undergo stage cracking to obtain the desired target products.

It is these patents that constitute the basic patents of the catalytic cracking process (MIP) that produces more isoparaffins, and have been widely used. At present, they have been applied to nearly 50 sets of catalytic cracking units, and have achieved huge economic and social benefits. Although the existing technology can obtain liquefied gas rich in isobutane and gasoline rich in isoparaffins, the processing of high-quality catalytic cracking feedstock oil, especially hydrogenated wax oil, results in low olefin content in gasoline, and the olefins in the liquefied gas The isobutene content is low, the product distribution is not optimized enough, and the petroleum resources are not fully utilized.

CN101362960A thinks that the single-ring aromatics content of diesel oil with different distillation ranges is different, and generally the single-ring aromatics content gradually decreases with the increase of the distillation range. In fact, dicyclic and tricyclic polycyclic aromatic hydrocarbons in diesel fraction, re-refined oil fraction and slurry fraction increased with the increase of distillation range. Moreover, the hydrogenation of aromatics is limited by thermodynamics. The equilibrium constant of hydrogenation of the first ring of fused-ring aromatics is relatively large, followed by that of the second ring. The equilibrium constant gradually decreases with the increase of the number of hydrogenated rings. From the perspective of reaction kinetics, the hydrogenation saturation rate of the first ring of polycyclic aromatic hydrocarbons is faster, and the hydrogenation saturation rate gradually decreases with the increase of the number of rings. For the heavy fraction of diesel oil rich in bicyclic aromatics, if one of the aromatic rings is hydrogenated, it can become a good catalytic cracking raw material for the production of high-octane gasoline, that is, one of the rings can be opened by catalytic cracking reaction The way of ring and side bond cleavage cracks the hydrocarbons in the diesel fraction into hydrocarbons in the gasoline fraction to increase the gasoline production. The most important thing is that the aromatics in the gasoline fraction have the characteristics of high research octane number and motor octane number, so the purpose of increasing gasoline research octane number and motor octane number can be achieved.

Contents of the invention

The purpose of the present invention is to provide a catalytic conversion method for increasing gasoline octane number barrel, so as to increase gasoline yield and its octane number simultaneously.

The method provided by the invention is: high-quality raw material oil and a thermally regenerated catalyst with low activity are contacted at the lower part of the reactor to undergo a cracking reaction, and the generated oil gas and the carbon-containing catalyst go up to undergo a selective hydrogen transfer reaction under a certain reaction environment And isomerization reaction, separation of reaction products, recycled catalyst after stripping and regeneration; cutting the diesel fraction in the reaction product into light diesel oil fraction and heavy diesel oil fraction, heavy diesel oil fraction alone or/and back-refined oil fraction, The oil slurry fraction is hydrotreated to obtain a hydrogenated heavy diesel oil fraction, and the hydrogenated heavy diesel oil fraction or/and the light oil oil fraction are returned to the reactor or/and other reactors for further reaction.

The method provided by the present invention is carried out in a reactor including realizing two types of different reactions, and the reactor is selected from one of equal-diameter riser, constant-linear-velocity riser, variable-diameter riser, and fluidized bed. It can be a composite reactor composed of equal diameter riser and fluidized bed.

Method provided by the invention is implemented like this:

(1) The preheated high-quality raw material oil enters the reactor and contacts with a thermally regenerated catalyst with an activity of 35 to 55, preferably 40 to 50, at a reaction temperature of 490°C to 620°C, preferably 500°C to 600°C, and a reaction time of 0.5 seconds to 2.0 The second is preferably 0.8 seconds to 1.5 seconds, and the cracking reaction occurs under the condition that the weight ratio of the catalyst to the raw oil (hereinafter referred to as the agent-oil ratio) is 3 to 15:1, preferably 3 to 12:1;

(2), the generated oil gas and the used catalyst go up, and selectivity occurs under the conditions of a reaction temperature of 420°C to 550°C, preferably 460°C to 500°C, and a reaction time of 2 seconds to 30 seconds, preferably 3 seconds to 15 seconds. Hydrogen transfer and isomerization reactions;

(3) Separating the reaction products to obtain liquefied gas rich in isobutene and gasoline and other products with moderate olefin content, the raw catalyst is stripped into the regenerator to be burnt and regenerated, and then recycled;

(4) The diesel fraction in the reaction product is cut into light diesel oil fraction and heavy diesel oil fraction, and the heavy diesel oil fraction or/and back-refined oil fraction and oil slurry fraction are subjected to hydrogenation treatment to obtain hydrogenated heavy diesel oil fraction, and the hydrogenated heavy diesel oil fraction is separated Or/and the light oil fraction returns to this reactor or/and other reactors for further reaction.

The cutting process of the diesel light fraction and the heavy fraction can be carried out separately or directly in the fractionation tower.

The temperature at the 5% point in the distillation range of the heavy fraction of diesel oil is not greater than 250°C.

The temperature at the 5% point in the distillation range of the diesel heavy fraction is not higher than 290°C.

The hydrogenated heavy diesel oil fraction or/and the light oil oil fraction can be returned to the device without restriction, for example, it can be returned to the bottom, middle and upper part of the reactor, and can be injected into the reactor alone or mixed with other raw materials, preferably no later than the raw material Oil is injected into the reactor system.

The diesel heavy fraction, re-refined oil fraction and oil slurry fraction entering the hydrotreating unit can be partly or completely entering, or fractions from other units.

The diesel heavy fraction can be hydrotreated alone, or it can be treated together with other raw materials.

The pressures of the cracking reaction in step (1), the hydrogen transfer reaction and the isomerization reaction in step (2) are all 130kPa～450kPa, and the weight ratio of water vapor to raw oil (hereinafter referred to as the water-oil ratio) is 0.03～0.3 :1, preferably 0.05 to 0.3:1.

The reactor suitable for this method can be selected from one of equal-diameter riser, equal-linear-velocity riser, fluidized bed or variable-diameter riser, or a composite reactor composed of equal-diameter riser and fluidized bed. .

The method provided by the invention can be carried out in equal-diameter risers, equal-linear-velocity risers or fluidized-bed reactors, wherein the medium-diameter risers are the same as the conventional catalytic cracking reactors in refineries, and the flow rate of fluid in the equal-linear-velocity risers Line speed is basically the same. Equal-diameter riser and constant-linear-velocity riser reactors are pre-lift section, first reaction zone, and second reaction zone from bottom to top; fluidized bed reactors are first reaction zone and second reaction zone from bottom to top , The height ratio of the first reaction zone and the second reaction zone is 10-40:90-60. When using a riser of equal diameter, a riser of constant linear velocity or a fluidized bed reactor, one or more cooling medium inlets are provided at the bottom of the second reaction zone, and/or a heat extractor is arranged in the second reaction zone, The height of the heat extractor 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 agent is one or a mixture of more than one selected from chilling agent, cooled regenerated catalyst and cooled semi-regenerated catalyst. Wherein the cold shock agent is selected from liquefied gas, crude gasoline, stable gasoline, diesel oil, heavy diesel oil or a mixture of more than one in any proportion in water; the cooled regenerated catalyst and the cooled semi-regenerated catalyst are the catalysts to be regenerated respectively After two-stage regeneration and one-stage regeneration and cooling, the carbon content of the regenerated catalyst is less than 0.1% by weight, preferably less than 0.05% by weight, and the carbon content of the semi-regenerated catalyst is 0.1% to 0.9% by weight, preferably 0.15% by weight. % by weight to 0.7% by weight.

The method provided by the present invention can also be carried out in the compound reactor that is made of equal-diameter riser and fluidized bed, the equal-diameter riser of the bottom is the first reaction zone, and the fluidized bed of the top is the second reaction zone, respectively controlled The temperature and reaction time of each reaction zone. One or more cooling shock 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% to 90% of the height of the second reaction zone. The temperature and reaction time of each reaction zone are controlled separately. The chilling agent is one or a mixture of more than one selected from chilling agent, cooled regenerated catalyst and cooled semi-regenerated catalyst. Wherein the cold shock agent is selected from liquefied gas, crude gasoline, stable gasoline, diesel oil, heavy diesel oil or a mixture of more than one in any proportion in water; the cooled regenerated catalyst and the cooled semi-regenerated catalyst are the catalysts to be regenerated respectively After two-stage regeneration and one-stage regeneration and cooling, the carbon content of the regenerated catalyst is less than 0.1% by weight, preferably less than 0.05% by weight, and the carbon content of the semi-regenerated catalyst is 0.1% to 0.9% by weight, preferably 0.15% by weight. % by weight to 0.7% by weight.

The method provided by the present invention can also be carried out in a variable-diameter riser reactor (referring to ZL99105903.4), and the structural characteristics of the reactor are as shown in Figure 1: the riser reactor is successively the same as each other along the vertical direction from bottom to top. The pre-lift section a of the shaft, the first reaction zone b, the second reaction zone c with enlarged diameter, the outlet zone d with reduced diameter, and a section of horizontal pipe e is connected at the end of the outlet zone. The junction of the first and second reaction zones is frustum-shaped, and the apex angle α of the isosceles trapezoid in its longitudinal section is 30° to 80°; The base angle β is 45°～85°.

The sum of the heights of the pre-lift section, the first reaction zone, the second reaction zone and the outlet zone of the reactor is the total height of the reactor, which is generally 10-60 meters.

The diameter of the pre-lift section is the same as that of a conventional equal-diameter riser reactor, generally 0.02 to 5 meters, and its height accounts for 5% to 10% of the total height of the reactor. The role of the pre-lift section is to move and accelerate the regenerated catalyst upwards in the presence of a pre-lift medium, the same as that used in conventional equal-diameter riser reactors, selected from steam or dry gas.

The structure of the first reaction zone is similar to a conventional equal-diameter riser reactor, and its diameter can be the same as that of the pre-lift section, or slightly larger than that of the pre-lift section. The ratio of the diameter of the first reaction zone to the diameter of the pre-lift section is 1.0～2.0:1, its height accounts for 10%～30% of the total height of the reactor. After the feedstock oil and catalyst are mixed in this zone, the cracking reaction mainly occurs at a relatively high reaction temperature, catalyst-to-oil ratio, and short residence time (generally 0.5 seconds to 2.5 seconds).

The second reaction zone is thicker than the first reaction zone, the ratio of its diameter to the diameter of the first reaction zone is 1.5-5.0:1, and its height accounts for 30%-60% of the total height of the reactor. Its function is to reduce the flow rate and reaction temperature of oil gas and catalyst. The method for reducing the reaction temperature in this zone can be to inject a cooling shock medium from the junction of this zone and the first reaction zone, and/or by setting a heat collector in this zone to take away part of the heat to reduce the reaction temperature in this zone, so as to achieve The purpose of inhibiting secondary cracking reaction, increasing isomerization reaction and hydrogen transfer reaction. The chilling agent is one or a mixture of more than one selected from chilling agent, cooled regenerated catalyst and cooled semi-regenerated catalyst. Wherein the cold shock agent is selected from liquefied gas, crude gasoline, stable gasoline, diesel oil, heavy diesel oil or a mixture of more than one in any proportion in water; the cooled regenerated catalyst and the cooled semi-regenerated catalyst are the catalysts to be regenerated respectively After two-stage regeneration and one-stage regeneration and cooling, the carbon content of the regenerated catalyst is less than 0.1% by weight, preferably less than 0.05% by weight, and the carbon content of the semi-regenerated catalyst is 0.1% to 0.9% by weight, preferably 0.15% by weight. % by weight to 0.7% by weight. If a heat extractor is provided, its height accounts for 50% to 90% of the height of the second reaction zone. The residence time of the stream in the reaction zone can be longer, ranging from 2 seconds to 30 seconds.

The structure of the outlet zone is similar to the top outlet part of a conventional equal-diameter riser reactor, the ratio of its diameter to the diameter of the first reaction zone is 0.8-1.5:1, and its height accounts for 0-20% of the total height of the reactor. The stream can stay in this zone for a certain period of time to suppress the overcracking reaction and thermal cracking reaction and increase the fluid flow rate.

One end of the horizontal pipe is connected to the outlet area, and the other end is connected to the settler; when the height of the outlet area is 0, that is, when the riser reactor has no outlet area, one end of the horizontal pipe is connected to the second reaction area, and the other end is connected to the settler . The function of the horizontal pipe is to transport the reaction product and the catalyst to the separation system for gas-solid separation. Its diameter is determined by those skilled in the art according to specific conditions. The function of the pre-lift section is to lift the regenerated catalyst into the first reaction zone in the presence of the pre-lift medium.

The high-quality raw material oils suitable for this method can be petroleum distillates with different boiling ranges. Specifically, the high-quality raw material oil is selected from one or more of atmospheric overhead oil, gasoline, catalytic gasoline, diesel oil, straight-run wax oil, and hydrogenated wax oil.

The two reaction zones in the method can be applied to all catalysts of the same type, which can be amorphous silica-alumina catalysts or zeolite catalysts. The active components of zeolite catalysts are selected from Y-type zeolite, HY-type zeolite, ultra-stable Y-type Type zeolite, ZSM-5 series zeolite or high silica zeolite with five-membered ring structure, ferrierite zeolite or a mixture of more than one in any proportion, the zeolite may contain rare earth and/or phosphorus, and may not contain rare earths and phosphorus.

The two reaction zones in the method may also be suitable for different types of catalysts, and the different types of catalysts may be catalysts with different particle sizes and/or catalysts with different apparent packing densities. Catalysts with different particle sizes and/or active components on catalysts with different apparent bulk densities are selected from different types of zeolites. A mixture of one or more than one of ferrierite and ferrierite with a membered ring structure in any proportion. The zeolite may contain rare earth and/or phosphorus, or may not contain rare earth and phosphorus. Catalysts with different particle sizes and/or high and low apparent bulk densities can enter different reaction zones, for example, catalysts with large particles containing ultrastable Y-type zeolite enter the first reaction zone to increase cracking reactions, and rare earth Y-type zeolite containing Small particles of zeolite catalysts enter the second reaction zone to increase the hydrogen transfer reaction. Catalysts with different particle sizes are stripped in the same stripper and regenerated in the same regenerator, and then the large and small particle catalysts are separated, and the small particle catalysts are cooled. into the second reaction zone. Catalysts with different particle sizes are separated by 30-40 microns, and catalysts with different apparent bulk densities are separated by 0.6-0.7 g/cm ³ .

The catalyst with lower activity applicable to the method refers to a catalyst with an activity of 35-55, preferably 40-50. It can be measured by measuring methods in the prior art: enterprise standard RIPP 92-90--micro-reactive test method of catalytic cracking "Petrochemical Analysis Method (RIPP Test Method)", Yang Cuiding et al., 1990, hereinafter referred to as RIPP 92-90. Described catalyst activity is represented by light oil slight reaction activity (MA), and its calculation formula is MA=(gasoline output+gas output+coke output lower than 204 ℃ in the product)/feed total amount*100%=in the product Gasoline yield + gas yield + coke yield below 204°C. The evaluation conditions of the light oil micro-reaction device (refer to RIPP 92-90) are: the catalyst is broken into particles with a diameter of 420-841 microns, the loading capacity is 5 grams, and the reaction raw material is straight-run light oil with a distillation range of 235-337 ° C. For diesel, the reaction temperature is 460°C, the weight space velocity is 16 hours ^-1 , and the agent-oil ratio is 3.2.

The catalyst used in the hydrogenation treatment can be a VIB group metal and/or a VIII group metal catalyst supported on alumina and/or amorphous silica-alumina carrier, and the preferred hydrogenation treatment catalyst is composed of 0 to 10% by weight additive , 1 to 9 weight percent of one or more metals of group VIII, 12 to 39 weight percent of one or more metals of group VIB and the balance of alumina and/or amorphous silica-alumina carrier, wherein the The above-mentioned additives are selected from non-metal elements and metal elements such as fluorine, phosphorus, titanium, and platinum. The VIB group metal is selected from Mo or/and W, and the VIII group metal is selected from Co or/and Ni.

The process conditions of the hydrogenation treatment are: hydrogen partial pressure 1.0-20.0MPa, reaction temperature 250-450°C, volume space velocity 0.1-20h ^-1 .

The hydrotreating can be performed in a fixed fluidized bed, or in other ways.

The advantages of the present invention are:

1. If adopting conventional equal-diameter riser or fluidized bed reactor to implement the present invention, only need to reduce the processing capacity and prolong the reaction time can be implemented.

2. If a variable-diameter riser reactor is used, the advantage of this reactor is that it not only retains the higher reaction temperature and solvent-oil ratio at the bottom of the conventional riser reactor to increase the primary cracking reaction, but also suppresses the overcracking and thermal cracking reactions at the top , and prolong the reaction time at a lower reaction temperature in the upper part of the reactor to increase the isomerization reaction and hydrogen transfer reaction of olefins.

3. Use the method provided by the invention to increase the gasoline yield by more than 5% by weight, increase the octane number of the gasoline research method by 1 unit, and increase the octane number of the motor method by 1 unit.

Description of drawings

Figure 1 is a schematic diagram of a new riser reactor, where a, b, c, d, and e in the figure represent the pre-lift section, the first reaction zone, the second reaction zone, the outlet zone, and the horizontal pipe, respectively.

Fig. 2 is a schematic flow chart of the best embodiment of the present invention. The descriptions of each number in the accompanying drawings are as follows:

1, 3, 4, 6, 11, 13, 17, 18, 20, 21, 22, 23, 24, 25, 26, 28, 29, 30, 31, 32, 33 all represent pipelines; 2 is the riser 5 and 7 are the first reaction zone and the second reaction zone of the riser respectively; 8 is the outlet zone of the riser; 9 is the settler, 10 is the cyclone separator, 12 is the stripper, and 14 is the waiting Raw inclined pipe, 15 is a regenerator, 16 is a regeneration inclined pipe, 19 is a separation system, and 27 is a hydrotreating unit.

Detailed ways

The present invention has different embodiments, one of the embodiments:

At the bottom of a conventional equal-diameter riser reactor, the preheated feed oil contacts the hot regenerated catalyst with low activity for a cracking reaction, and the generated oil gas and spent catalyst travel upward to contact the injected cooled regenerated catalyst, followed by heterogeneity. Structured reaction and hydrogen transfer reaction, the effluent after the reaction enters the settler; the reaction product is separated, and the raw catalyst is stripped and regenerated and divided into two parts, one part enters the bottom of the reactor, and the other part enters the reaction after cooling The lower part of the device. The diesel fraction in the reaction product is cut into light diesel oil fraction and heavy diesel oil fraction, and the heavy diesel oil fraction is separately or/and re-refined oil fraction and oil slurry fraction are subjected to hydrotreating to obtain hydrogenated heavy diesel oil fraction, and the hydrogenated heavy diesel oil fraction is alone or/ And light oil fraction returns to this reactor or/and other reactors for further reaction.

The second implementation mode:

At the bottom of a conventional equal-diameter riser reactor, the preheated feedstock oil contacts the hot regenerated catalyst with low activity for a cracking reaction, and the generated oil gas and spent catalyst ascend to contact with the injected chiller and cooled semi-regenerated catalyst , followed by isomerization reaction and hydrogen transfer reaction, the effluent after the reaction enters the settler; the reaction product is separated, and the raw catalyst is stripped, then enters the two-stage regenerator for charring, and comes out of the first-stage regenerator The semi-regenerated catalyst enters the middle and lower part of the reactor after cooling down, and the regenerated catalyst coming out of the second stage regenerator directly returns to the bottom of the reactor without cooling down. The diesel fraction in the reaction product is cut into light diesel oil fraction and heavy diesel oil fraction, and the heavy diesel oil fraction is separately or/and re-refined oil fraction and oil slurry fraction are subjected to hydrotreating to obtain hydrogenated heavy diesel oil fraction, and the hydrogenated heavy diesel oil fraction is alone or/ And light oil fraction returns to this reactor or/and other reactors for further reaction.

The third implementation mode:

For a catalytic cracking unit with a conventional riser-fluidized bed reactor, the preheated conventional cracking feedstock enters from the lower part of the riser to contact with the thermally regenerated catalyst with low activity, and the oil and gas generated after the reaction go up to the top of the riser , contact with the cooled catalyst to continue the reaction, and the effluent after the reaction enters the settler; the reaction product is separated, and the raw catalyst is stripped and regenerated and divided into two parts, one of which enters the lower part of the riser, and the other part after cooling into the top of the riser. The diesel fraction in the reaction product is cut into light diesel oil fraction and heavy diesel oil fraction, and the heavy diesel oil fraction is separately or/and re-refined oil fraction and oil slurry fraction are subjected to hydrotreating to obtain hydrogenated heavy diesel oil fraction, and the hydrogenated heavy diesel oil fraction is alone or/ And light oil fraction returns to this reactor or/and other reactors for further reaction.

The fourth implementation mode:

This embodiment is the best embodiment of the present invention.

For the catalytic cracking unit with a new type of variable-diameter riser reactor, the preheated conventional cracking raw material enters from the lower part of the first reaction zone of the reactor to contact with the thermally regenerated catalyst with low activity, and a cracking reaction occurs, and the oil gas generated after the reaction Go up to the lower part of the second reaction zone of the reactor and contact with the cooled catalyst to carry out hydrogen transfer reaction and isomerization reaction. After the reaction, the effluent enters the settler; the reaction product is separated, and the raw catalyst is stripped, regenerated and then enters the second lower part of the reaction zone. The diesel fraction in the reaction product is cut into light diesel oil fraction and heavy diesel oil fraction, and the heavy diesel oil fraction is separately or/and re-refined oil fraction and oil slurry fraction are subjected to hydrotreating to obtain hydrogenated heavy diesel oil fraction, and the hydrogenated heavy diesel oil fraction is alone or/ And light oil fraction returns to this reactor or/and other reactors for further reaction.

The method provided by the present invention is not limited thereto.

The method provided by the present invention will be further described below in conjunction with the accompanying drawings, but the present invention is not limited thereto.

Figure 2 is the flow chart of the catalytic conversion method for increasing the content of isobutene and gasoline olefins in liquefied gas by using a riser reactor with variable diameter. The shape and size of equipment and pipelines are not limited by the accompanying drawings, but determined according to specific conditions.

The pre-lift steam enters from the pre-lift section 2 of the riser through the pipeline 1, and the hot regenerated catalyst enters the pre-lift section of the riser through the regenerated inclined pipe 16 and is lifted by the pre-lift steam. The preheated raw oil enters from the pre-lifting section of the riser through the pipeline 4 and the atomized steam from the pipeline 3 in a certain proportion, mixes with the hot catalyst and enters the first reaction zone 5, and undergoes cracking reaction under certain conditions. The reactant stream is mixed with the chiller and/or cooled catalyzer (not shown among the figures) from the pipeline 6 and enters the second reaction zone 7 for secondary reaction. The reacted stream enters the outlet zone 8, and the reaction zone increases the flow rate. The linear velocity makes the reactant flow quickly enter the settler 9 and the cyclone separator 10 in the gas-solid separation system, and the reaction product is separated through the pipeline 11 to the separation system 19, and the obtained dry gas is drawn out through the pipeline 20, and the liquefied gas is drawn out through the pipeline 21. The gasoline fraction is drawn through the pipeline 22, the diesel light fraction is drawn through the pipeline 23, the diesel heavy fraction is introduced into the hydrotreating unit 27 through the pipeline 24, the back-refined oil fraction is drawn through the pipeline 25, or/and introduced into the hydrotreating unit 27, and the oil slurry fraction is drawn through the pipeline 25 The pipeline 26 leads out, or/and leads into a hydrotreating unit 27 . After the reaction, the raw catalyst with charcoal enters the stripper 12, and after being stripped by the water vapor from the pipeline 13, it enters the regenerator 15 through the inclined pipe 14, and the raw catalyst is burnt and regenerated in the air from the pipeline 17. The gas exits the regenerator through the pipeline 18, and the hot regenerated catalyst returns to the bottom of the riser through the regenerating inclined pipe 16 for recycling. The hydrotreated heavy diesel oil fraction is injected into the riser through the pipeline 28 and the pipeline 29 and the atomized steam from the pipeline 30 in a certain proportion, or/and is injected into the riser through the pipeline 32 and the atomized steam from the pipeline 33 in a certain proportion, or/ And through pipeline 31 and pipeline 4, it is injected into the riser together with raw oil.

The following examples will further illustrate the present invention, but do not limit the present invention thereby. The properties of the feed oil and catalyst used in the examples and comparative examples are listed in Table 1 and Table 2 respectively. The catalysts in Table 2 are all produced by Qilu Catalyst Factory of China Petrochemical Corporation. The ZCM-7 catalyst in Table 2 was aged at 800°C and 100% steam for 30 hours to obtain ZCM-7 with an activity level of 45; similarly, the CGP-1 catalyst in Table 2 was aged at 800°C and 100% steam for 30 hours hours, an activity level of 50 CGP-1 was obtained.

The hydrotreating unit used is a fixed fluidized bed unit, and the hydrotreating catalyst is RN-10 catalyst, which is produced by Changling Catalyst Factory of China Petrochemical Corporation.

Example 1

This example illustrates the use of the method provided by the present invention and the use of catalysts with different activity levels to increase the isobutene content in liquefied gas and the olefin content in gasoline in a medium-sized variable-diameter riser reactor.

The total height of the pre-lift section, the first reaction zone, the second reaction zone and the outlet zone of the reactor is 15 meters, the diameter of the pre-lift section is 0.025 meters, and its height is 1.5 meters; the diameter of the first reaction zone is 0.025 meters, and its height It is 4 meters; the diameter of the second reaction zone is 0.1 meters, and its height is 6.5 meters; the diameter of the outlet zone is 0.025 meters, and its height is 3 meters; is 45°; the base angle of the isosceles trapezoid in the longitudinal section of the junction of the second reaction zone and the outlet zone is 60°.

The feed oil A listed in preheated table 1 enters the reactor, and in the presence of water vapor, contacts with the catalyst ZCM-7 listed in hot table 2. The catalyst activity of ZCM-7 is 45, and the reaction product is separated to obtain For liquefied gas, gasoline and other products, the 5% point of the heavy fraction of diesel oil is 260°C, and the heavy fraction of diesel oil is sent to the hydrotreating unit for treatment. The reaction temperature of the hydrotreating unit is 345°C, the hydrogen partial pressure is 2.5MPa, and the volume space velocity is 1.3h ^-1 , the hydrotreated diesel heavy fraction is injected into the riser through the pipeline 29 in Fig. 2, the amount of atomized water vapor is 10% (accounting for the hydrotreated diesel heavy fraction), and the catalyst to be regenerated is stripped and enters the regenerator. The regenerated catalyst is recycled after being burnt.

The operating conditions, product distribution and gasoline properties tested are listed in Table 3.

Comparative example 1

The reactor type and operating conditions are exactly the same as in Example 1, the feed oil used is also the feed oil A listed in table 1, the catalyst is also the catalyst ZCM-7 listed in table 2, and the catalyst activity is 45, but at this time the diesel fraction It is directly exported as a product without cutting. The operating conditions, product distribution and gasoline properties tested are listed in Table 3.

As can be seen from Table 3, compared with the re-cracking mode without hydrotreating of heavy fractions of diesel oil, when the mode of hydrotreating and re-cracking of heavy fractions of diesel oil (pattern of embodiment 1) is adopted, the gasoline yield increases by 5.4 units, and the research method octane Value (RON) was increased by 0.4 units, Motor Octane Number (MON) was increased by 0.5 units, Research Octane Number Barrel (Research Octane Number multiplied by gasoline yield divided by 100) was increased from 50.1 to 55.2, an increase of 5.1 units, a whopping 10.3 percent increase.

Example 2

This example illustrates the use of the method provided by the present invention and the use of catalysts with different activity levels to increase the isobutene content in liquefied gas and the olefin content in gasoline in a medium-sized variable-diameter riser reactor.

The total height of the pre-lift section, the first reaction zone, the second reaction zone and the outlet zone of the reactor is 15 meters, the diameter of the pre-lift section is 0.025 meters, and its height is 1.5 meters; the diameter of the first reaction zone is 0.025 meters, and its height It is 4 meters; the diameter of the second reaction zone is 0.1 meters, and its height is 6.5 meters; the diameter of the outlet zone is 0.025 meters, and its height is 3 meters; is 45°; the base angle of the isosceles trapezoid in the longitudinal section of the junction of the second reaction zone and the outlet zone is 60°.

The raw material oil A listed in the table 1 of preheating enters in this reactor, under the presence of water vapor, contacts with the catalyst CGP-1 listed in the table 2 of heat and reacts, and CGP-1 catalyst activity is 50, and separation reaction product obtains For liquefied gas, gasoline and other products, the 5% point of diesel heavy fraction is 250°C, and the diesel heavy fraction is sent to the hydrotreating unit for treatment. The reaction temperature of the hydrotreating unit is 340°C, the hydrogen partial pressure is 2.8MPa, and the volume space velocity is 1.2h ^-1 , the heavy fraction of hydrotreated diesel oil is injected into the riser through the pipeline 29 in Fig. 2, the amount of atomized water vapor is 10% (accounting for the heavy fraction of hydrotreated diesel oil), and the catalyst to be regenerated is stripped and enters the regenerator. The regenerated catalyst is recycled after being burnt.

The operating conditions, product distribution and gasoline properties tested are listed in Table 4.

Comparative example 2

The reactor type and operating conditions are exactly the same as in Example 2, the feed oil used is also the feed oil A listed in table 1, the catalyst is also the catalyst CGP-1 listed in table 2, and the catalyst activity is 50, but the diesel fraction is now It is directly exported as a product without cutting. The operating conditions, product distribution and gasoline properties tested are listed in Table 4.

As can be seen from Table 4, compared with the re-cracking mode without hydrotreating of heavy fractions of diesel oil, when the mode of hydrotreating and re-cracking of heavy fractions of diesel oil is adopted (pattern of embodiment 2), the gasoline yield increases by 6.0 units, and RON increases by 1.0 units, MON increased by 1.1 units, and the research octane barrel increased from 40.1 to 46.2, an increase of 6.1 units, an increase of 15.2%.

Table 1

Raw oil number A Raw oil name Hydrogenated wax oil Density (20℃), kg/ ^m3 899.3 Kinematic viscosity, ^mm2 /s 80℃ 16.22 100℃ 9.29 Carbon residue, wt% 0.30 Freezing point, ℃ 44 Basic nitrogen, ppm 293 Total nitrogen, wt% 0.08 Sulfur, wt% 0.12 Carbon, weight % 87.01 Hydrogen, weight % 12.85 Distillation range, ℃ initial boiling point 284 10% 394 30% 433 50% 463 70% 495 90% / end point /

Table 2

Catalyst number A B brand name ZCM-7 CGP-1 Zeolite type USY REY-USY-ZRP Chemical composition, wt% Alumina 46.4 52.0 sodium oxide 0.22 0.14 iron oxide 0.32 0.30 Apparent density, kg/ ^m3 600 740 Pore volume, ml/g 0.32 0.37 Specific surface area, ^m2 /g 217 263 Sieve composition, wt% 0～40 microns 16.1 20.3 40～80 microns 54.1 / ＞80 microns 29.8 /

table 3

Example 1 Comparative example 1 ZCM-7 catalyst activity 45 45 Reaction temperature, ℃ The first reaction zone 550 550 Second reaction zone 500 500 Dwell time, seconds 5.5 5.5 The first reaction zone 2.0 2.0 Second reaction zone 3.5 3.5 Agent oil ratio 5.0 5.0 Water to oil ratio 0.1 0.1 Product distribution, weight % dry gas 1.5 1.4 Liquefied gas 18.0 17.3 Of which isobutylene 2.0 2.0 gasoline 60.4 55.0 diesel fuel 10.9 17.8 heavy oil 6.4 6.0 Coke 2.9 2.5 Total 100.1 100.0 Liquid yield, weight % 89.3 90.1 octane number RON 91.4 91.0 MON 81.2 80.7 Research Octane Bucket 55.2 50.1 Distillation range, ℃ Initial boiling point to dry point 38～200 38～200 Family composition, weight % alkanes 38.8 40.5 Naphthenic 7.8 7.3 Olefins 27.9 29.3 Aromatics 25.5 22.9

Table 4

Example 2 Comparative example 2 CGP-1 catalyst activity 50 50 Reaction temperature, ℃ The first reaction zone 550 550 Second reaction zone 505 505 Agent oil ratio 6.0 6.0 Response time, seconds 6.0 6.0 Among them, the first reaction zone 1.3 1.3 Second reaction zone 4.7 4.7 Water to oil ratio 0.1 0.1 Product distribution, weight % dry gas 1.9 1.8 Liquefied gas 29.3 28.5 Of which propylene 10.8 10.6 Isobutylene 4.1 4.1 gasoline 48.9 42.9 diesel fuel 10.4 18.0 heavy oil 6.8 6.5 Coke 2.8 2.3 Liquid yield, weight % 88.6 89.4 gasoline octane number RON 94.5 93.5 MON 82.6 81.5 Research Octane Bucket 46.2 40.1 Distillation range, ℃ Initial boiling point to dry point 38～200 38～200 Gasoline family composition, weight % alkanes 33.6 35.9 Naphthenic 7.8 7.6 Olefins 26.0 27.9 Aromatics 32.6 28.6

Claims (12)

1. one kind is improved the catalysis conversion method of gasoline octane rating bucket, it is characterized in that high quality raw material oil and the lower contacts generation cracking reaction of active lower hot regenerated catalyst at reactor, the oil gas generated and under certain reaction environment, optionally hydrogen transfer reactions and isomerization reaction occur containing the catalyzer of charcoal is up, reaction product isolated, reclaimable catalyst uses through stripping, regeneration Posterior circle; Diesel oil distillate in reaction product is cut into solar oil cut and heavy gas oil cut, heavy gas oil cut is separately or/and recycle stock cut, slurry oil cut carry out hydrotreatment obtain hydrogenation heavy gas oil cut, and hydrogenation heavy gas oil cut is separately or/and solar oil cut returns this reactor or/and other reactors react further; The hot regenerated catalyst activity that described activity is lower is 35 ~ 55; Cracking reaction condition is: temperature of reaction 490 DEG C ~ 620 DEG C, 0.5 second ~ 2.0 seconds reaction times, the weight ratio 3 ~ 15: 1 of catalyzer and stock oil; Hydrogen transfer reactions and isomerization reaction condition are: temperature of reaction 420 DEG C ~ 550 DEG C, and the reaction times is 2 seconds ~ 30 seconds.

2., according to the method for claim 1, it is characterized in that one or more in atmospheric overhead, gasoline, diesel oil, straight-run gas oil, hydrogenation wax oil of described high quality raw material grease separation.

3., according to the method for claim 2, it is characterized in that described gasoline is catalytic gasoline.

4., according to the method for claim 1, it is characterized in that the hot regenerated catalyst activity that described activity is lower is 40 ~ 50.

5., according to the method for claim 1, it is characterized in that described cracking reaction condition is: temperature of reaction 500 DEG C ~ 600 DEG C, 0.8 second ~ 1.5 seconds reaction times, the weight ratio 3 ~ 12: 1 of catalyzer and stock oil.

6., according to the method for claim 1, it is characterized in that described hydrogen transfer reactions and isomerization reaction condition are: temperature of reaction 460 DEG C ~ 500 DEG C, the reaction times is 3 seconds ~ 15 seconds.

7., according to the method for claim 1, it is characterized in that described cracking reaction, the pressure of hydrogen transfer reactions and isomerization reaction is 130kPa ~ 450kPa, the weight ratio of water vapor and stock oil is 0.03 ~ 0.3: 1.

8. according to the method for claim 1, it is characterized in that described reactor be selected from equal diameter riser tube, etc. one of in linear speed riser tube, fluidized-bed or reducing riser tube, or the compound reactor be made up of equal diameter riser tube and fluidized-bed.

9. according to the method for claim 8, it is characterized in that described reducing riser tube to be vertically followed successively by the outlet area of pre lift zone coaxial each other, the first reaction zone, the second reaction zone of enlarged-diameter, reduced from bottom to up, be connected with one section of level pipe at outlet area end, wherein the diameter of second reaction zone and the diameter ratio of the first reaction zone are 1.5 ~ 5.0: 1.

10., according to the method for claim 1, it is characterized in that the processing condition of described hydrotreatment are: hydrogen dividing potential drop 1.0 ~ 20.0MPa, temperature of reaction 250 ~ 450 DEG C, volume space velocity 0.1 ~ 20h ^-1.

11., according to the method for claim 1, is characterized in that in described diesel oil last running boiling range, 5% temperature is not more than 250 DEG C.

12., according to the method for claim 11, is characterized in that in described diesel oil last running boiling range, 5% temperature is not more than 290 DEG C.