CN110643380B

CN110643380B - Method for converting coal pyrolysis product into gasoline, diesel oil and hydrogen

Info

Publication number: CN110643380B
Application number: CN201910778005.XA
Authority: CN
Inventors: 万会军; 杨强; 许文静; 赵青青; 郭强; 田磊; 杨勇; 李永旺
Original assignee: SYNEFUELS CHINA Inc; Synfuels China Technology Co Ltd
Current assignee: SYNEFUELS CHINA Inc.; Zhongke Synthetic Oil Technology Co Ltd
Priority date: 2019-08-22
Filing date: 2019-08-22
Publication date: 2020-10-13
Anticipated expiration: 2039-08-22
Also published as: CN110643380A

Abstract

The invention relates to a method for converting coal pyrolysis products into high-octane gasoline, diesel oil and hydrogen, wherein the coal pyrolysis products are firstly divided into hydrocarbon mixtures of five distillation sections according to the light weight of the distillation sections, and then products of each distillation section are treated by the combination of technologies such as tail gas treatment, catalytic cracking, deep hydrogenation, solvent oil hydrogenation, coking treatment, catalytic reforming, alkylation treatment, optional C5/C6 isomerization treatment and the like, so that the coal pyrolysis products are converted into the high-octane gasoline, the diesel oil and the hydrogen with high efficiency and good economic benefit.

Description

Method for converting coal pyrolysis product into gasoline, diesel oil and hydrogen

Technical Field

The invention belongs to the field of coal deep processing, and particularly relates to a method for converting coal pyrolysis products into gasoline, diesel oil and hydrogen.

Background

China has abundant coal resources, develops the coal-to-oil chemical industry, and has important significance for relieving the contradiction between supply and demand of petroleum and realizing clean utilization of coal.

There are three main ways to produce oil products by coal liquefaction: a partial liquefaction process in which a portion of the coal is converted into tar after dry distillation (also referred to as coal pyrolysis); the coal is completely converted into an oil product through direct hydrogenation in a direct liquefaction process; and the indirect liquefaction process of producing synthetic gas by gasifying coal and producing high-quality oil products by Fischer-Tropsch synthesis. The three approaches have already entered into industrial production as early as the last century and have become the major oil production approach in world war II in Germany.

With respect to the existing coal liquefaction technology, the biggest challenge of the large-scale dry distillation technology is to develop a high-efficiency and environment-friendly process technology and an effective treatment and utilization method of normal-pressure dry distillation gas (including flue gas and fuel gas), and although the indirect liquefaction process is verified in the aspects of large scale, safety, reliability and environmental protection, the biggest problem is that the consumption of oil and coal per ton is large due to the low calorific value of gasification raw materials and the like. The direct liquefaction process still has the problems of equipment manufacture, safe and stable operation, low diesel oil quality of products and the like.

CN101311246A discloses a mild hydrogenation upgrading method for direct coal liquefaction oil, which comprises the steps of mixing the direct coal liquefaction oil with hydrogen, feeding the mixture into a hydrogenation stabilization reactor, carrying out desulfurization, denitrification, dearomatization and olefin saturation reactions under the action of a hydrogenation stabilization catalyst, separating and fractionating reaction products to obtain naphtha fraction and diesel fraction, mixing the obtained diesel fraction with hydrogen, feeding the mixture into a hydrogenation modification reactor, carrying out contact reaction with a hydrogenation modification catalyst, and separating and fractionating reaction products to obtain a diesel product. The method provided by the invention effectively combines the stable hydrogenation unit and the hydrogenation modification unit together, and can produce the diesel oil product with the cetane number reaching 45.

CN102206511A discloses a method for producing diesel oil by mixing coal direct liquefaction oil and delayed coking heavy diesel oil fraction for hydro-upgrading.

CN103965961A discloses a process for obtaining naphtha fraction, medium distillate and heavy distillate from coal direct liquefaction oil through full fraction hydrotreating, wherein the naphtha fraction is subjected to catalytic reforming to obtain gasoline.

However, in the processing schemes for direct coal liquefaction products disclosed in the prior art, the schemes for mainly producing diesel oil can produce a large amount of naphtha and liquefied gas as by-products, and the product economy is poor, so that the economy of the schemes is poor. The scheme for producing gasoline mainly discloses that gasoline is produced by a certain fraction section in coal liquefaction products, the yield of the gasoline is low, and related components in the rest fractions are not effectively utilized.

Therefore, aiming at the problems in the prior art, the upgrading process of the coal pyrolysis product of the whole-distillation section is developed, and integration and improvement are carried out from the source, so that the utilization efficiency of the Fischer-Tropsch synthesis product can be improved, and the economic benefit can be improved.

Disclosure of Invention

The invention aims to solve the technical problems of low processing and utilization efficiency of coal pyrolysis products and poor economical efficiency of product processing routes in the prior art, and provides an improved and more economical complete oil product processing flow by optimizing and combining the prior art and improving the process according to the characteristics of the coal pyrolysis products, which is mainly used for preparing clean high-octane gasoline, diesel oil and hydrogen.

The inventor finds that the coal pyrolysis product can be converted into gasoline, diesel oil and hydrogen with high efficiency and good economic benefit by combining the technologies of tail gas treatment, catalytic cracking, deep hydrogenation, solvent oil hydrogenation, coking treatment, catalytic reforming, alkylation treatment, C5/C6 isomerization treatment and the like aiming at the existing coal pyrolysis product.

Accordingly, the present invention discloses a process for converting coal pyrolysis products into high octane gasoline, diesel and hydrogen, wherein the process comprises the steps of:

(1) dividing the coal pyrolysis product into hydrocarbon mixtures of five distillation sections of tail gas, pyrolysis light oil, pyrolysis heavy oil, solvent oil and bottom oil residue according to the weight of the distillation;

(2) washing and separating the tail gas to obtain heavy hydrocarbon and pretreated tail gas, performing absorption stabilization treatment on the heavy hydrocarbon to obtain absorption-stabilized naphtha and absorption-stabilized LPG, and performing MDEA desulfurization treatment and pressure swing adsorption treatment on the pretreated tail gas to obtain hydrogen;

(3) enabling the solvent oil to enter a solvent oil hydrogenation unit, and performing hydrogenation treatment to obtain hydrogen-supplying solvent oil, hydrogenated light oil and hydrogenated gas;

(4) enabling the bottom-reduced oil residue to enter a coking unit to obtain coking heavy oil and coking light oil;

(5) mixing the pyrolysis light oil, the coking light oil, the pyrolysis heavy oil, the coking heavy oil, the hydrogenation light oil and the hydrogenation gas, sending the mixture to a deep hydrogenation unit, and performing hydrogenation treatment and fractionation to obtain deep hydrogenation diesel oil, heavy naphtha and light naphtha;

(6) carrying out catalytic cracking on part or all of the deeply hydrogenated diesel oil to obtain catalytic heavy gasoline, catalytic light gasoline, mixed carbon three and mixed carbon four; carrying out etherification treatment on the catalytic light gasoline to obtain etherified gasoline;

(7) carrying out catalytic reforming on the heavy naphtha and part or all of the absorption-stabilized naphtha to obtain reformed gasoline, reformed LPG and reformed gas; optionally, subjecting the light naphtha to C5/C6 isomerization treatment to obtain isomerized gasoline and isomerized LPG;

(8) treating the mixed C4 in an MTBE unit to obtain MTBE and ether C4;

(9) subjecting said post-ethereal carbon four, said mixed carbon three, said absorption-stabilized LPG and said reformed LPG and optionally said isomerized LPG to alkylation treatment to obtain an alkylated gasoline and an alkylated LPG;

(10) blending the alkylated gasoline, the catalytic heavy gasoline, the etherified gasoline, the reformed gasoline, the MTBE and optionally the isomerized gasoline to obtain the high-octane gasoline.

The invention has the following characteristics:

(i) provides an integrated process flow for processing the full-fraction coal pyrolysis product, so that all components in the coal pyrolysis product can be effectively utilized;

(ii) the method adopts a combined scheme of water washing separation, MDEA (N-methyldiethanolamine) desulfurization, PSA (pressure swing adsorption) and absorption stabilization treatment, deeply separates and utilizes tail gas which is one of components of coal pyrolysis products, and fully recovers effective components such as naphtha, LPG, hydrogen and the like; further, by processing naphtha in the tail gas into reformed gasoline and LPG in the tail gas into alkylated gasoline, the liquid yield is increased, and the additional value is improved;

(iii) the heavy naphtha and the light naphtha obtained after deep hydrogenation are respectively converted into the reformed gasoline and the isomerized gasoline, so that the yield of the naphtha with poor economy is reduced, a clean gasoline product with high octane value is further obtained, and the economic benefit is improved;

(iv) deep hydrogenated diesel can be converted into high octane catalytic gasoline (including catalytic light gasoline and catalytic heavy gasoline) by using a catalytic cracking process; or the deep hydrogenation diesel can be used as final clean diesel for direct subsequent utilization; the catalytic light gasoline is further etherified to generate etherified gasoline, so that the octane number of the gasoline is further improved;

(v) LPG in the catalytic cracking gas is separated into mixed C4 and mixed C three, isobutene in the mixed C four is processed into a gasoline additive MTBE (methyl tert-butyl ether) with high added value, olefin in the mixed C three is processed into alkylated gasoline, and the utilization efficiency of the LPG is improved.

Drawings

FIG. 1 is a schematic diagram showing a process for producing high octane gasoline, diesel and hydrogen from coal pyrolysis products (wherein the coking unit employs a flexicoking process);

FIG. 2 is a schematic diagram showing a process for producing high octane gasoline, diesel and hydrogen from coal pyrolysis products (wherein the coking unit employs a delayed coking process);

fig. 3 is a schematic diagram showing a process for producing high octane gasoline, diesel and hydrogen from coal pyrolysis products (wherein the absorption stabilized naphtha from tail gas treatment is not subjected to deep hydrotreating and the alkylation treatment employs ionic liquid alkylation technology).

Detailed Description

The following exemplary embodiments are only for explaining the aspects of the present invention, and are not intended to limit the scope of protection of the present application in any way.

In the present invention, unless otherwise specified, the pyrolysis light oil is mainly a hydrocarbon mixture having a distillation range of 90 ℃ or more and less than 175 ℃.

In the present invention, unless otherwise specified, the pyrolysis heavy oil is mainly a hydrocarbon mixture having a distillation range of 175 to 450 ℃.

In the present invention, unless otherwise specified, the mineral spirits are predominantly hydrocarbon mixtures with a boiling range above 290 ℃ and below 460 ℃.

In the present invention, unless otherwise specified, the reduced bottoms are hydrocarbon mixtures having a boiling range above 460 ℃.

In the present invention, unless otherwise specified, the term "coker light oil" is primarily a hydrocarbon mixture having a boiling range above that of the C5 hydrocarbons and below 175 ℃.

In the present invention, unless otherwise specified, the term "coker heavy oil" is primarily a hydrocarbon mixture having a boiling range of 175 ℃ to 540 ℃.

In the present invention, the term "heavy naphtha" means C7-C10 hydrocarbon fractions unless otherwise specified.

In the present invention, the term "light naphtha" means C5-C6 hydrocarbon fractions unless otherwise specified.

In the present invention, unless otherwise specified, the term "high octane gasoline" means a blending component of alkylated gasoline, catalytic heavy gasoline, etherified gasoline, reformed gasoline, MTBE and optionally isomerized gasoline, wherein the "high octane" means RON of 85 or more.

In the present invention, the term "part or all" means a part (for example, more than 0% to less than 100%) of the subject modified by the term or the whole thereof, unless otherwise specified.

In one embodiment, the present invention relates to a process for converting coal pyrolysis products to high octane gasoline, diesel and hydrogen, wherein the process comprises the steps of:

(8) treating the mixed C4 in an MTBE unit to obtain MTBE and ether C4;

In a preferred embodiment, in the step (2), the off-gas is first washed by a water washing tower and then cooled (preferably to 5 ℃ to 15 ℃, further preferably to 6 ℃ to 9 ℃) to perform the water washing separation. Ammonia in the tail gas can be removed by washing in the water washing tower, and a small amount of water and heavy hydrocarbon carried in the tail gas can be further separated by cooling the tail gas after washing.

In a preferred embodiment, in the step (2), the absorption stabilization treatment is performed using a combination of an absorption desorption column and a stabilization column. Preferably, the theoretical plate number of the absorption and desorption tower is 10-20, preferably 15-18, the tower top temperature is 40-50 ℃, preferably 42-48 ℃, the tower bottom temperature is 120-130 ℃, preferably 124-128 ℃, and the operation pressure is 0.5-1 MPaG, preferably 0.6-0.9 MPaG; preferably, the theoretical plate number of the stabilizer is 20-30, preferably 22-28, the tower top temperature is 60-70 ℃, preferably 62-68 ℃, the tower bottom temperature is 190-200 ℃, preferably 192-195 ℃, and the operating pressure is 0.5-1.5 MPaG, preferably 0.9-1.0 MPaG.

In a preferred embodiment, in the step (3), the operation conditions of the solvent oil hydrogenation unit are as follows: solvent oil hydrogenation volume space velocity of 0.1h^-1～1h^-1Preferably 0.5h^-1～1h^-1(ii) a The volume ratio of hydrogen to oil is 500-1500, preferably 1000-1500.

In a preferred embodiment, in step (3), the hydrogen-donating solvent oil is returned to the coal hydropyrolysis unit.

In a preferred embodiment, in step (4), the coking unit is a flexicoking unit. In a further preferred embodiment, in the flexicoking unit, the flexicoking reactor is operated at a temperature of 400 ℃ to 600 ℃, preferably 480 ℃ to 550 ℃; the operating pressure is from 0MPaG to 0.2MPaG, preferably from 0.05MPaG to 0.15 MPaG. In a further preferred embodiment, the reduced bottoms sludge is passed to the flexicoking unit and treated to further obtain a heating value of less than 1200Kcal/Nm³Of the low heating value gas. In a further preferred embodiment, the method comprisesAnd mixing the low-calorific-value gas with the pretreated tail gas, and performing MDEA desulfurization treatment and pressure swing adsorption treatment to obtain hydrogen. In a further preferred embodiment, the conditions of the MDEA desulfurization treatment are: the number of theoretical plates of the desulfurizing tower is 10-20, preferably 10-15, the operating pressure is 0.5 MpaG-1.0 MPaG, preferably 0.6 MpaG-0.9 MPaG, the tower top temperature is 35-45 ℃, preferably 38-42 ℃, and the amine liquid concentration is 20-30 wt%, preferably 22-26 wt%; the theoretical plate number of the regeneration tower is 10-20, preferably 10-15, the operation temperature is 120-130 ℃, preferably 120-125 ℃, and the operation pressure is 0.10-0.20 MPaG, preferably 0.10-0.15 MPaG. In a further preferred embodiment, the pressure swing adsorption process conditions are: the operation temperature of the pressure swing adsorption tower is less than or equal to 40 ℃, the pressure during adsorption is 2.0MPaG to 2.5MPaG, preferably 2.3MPaG to 2.5MPaG, and the pressure during regeneration is 0.02MpaG to 0.07MpaG, preferably 0.04MpaG to 0.05 MpaG.

In another preferred embodiment, in step (4), the coking unit is a delayed coking unit. In a further preferred embodiment, in the delayed coking unit, the overhead pressure of the coke drum is from 0.15MPaG to 0.2MPaG, preferably from 0.17MPaG to 0.19 MPaG; the outlet temperature of the coke drum feed furnace is from 495 ℃ to 505 ℃, preferably from 498 ℃ to 502 ℃. In a further preferred embodiment, the reduced bottoms sludge is passed to the delayed coking unit and further processed to obtain coke. In a further preferred embodiment, the coke is gasified and purified, and then subjected to Fischer-Tropsch synthesis to produce Fischer-Tropsch oil products.

In another preferred embodiment, in step (4), the coking unit is a fluid coking unit.

In a preferred embodiment, in the step (5), a part of the absorption-stabilized naphtha is mixed with the pyrolysis light oil, the coking light oil, the pyrolysis heavy oil, the coking heavy oil, and the hydrogenation light oil and the hydrogenation gas and sent to a deep hydrogenation unit.

In a preferred embodiment, in step (5), the operating conditions of the deep hydrogenation unit are: depth addingThe space velocity of hydrogen volume is 0.1h^-1～1h^-1Preferably 0.3h^-1～0.7h^-1The volume ratio of hydrogen to oil is 500-1500, preferably 1000-1400.

In a preferred embodiment, in the step (5), after hydrotreating, fractionation may also be performed to obtain a hydrogen-rich tail gas. Preferably, the hydrogen-rich tail gas is recycled to the step (2) for water washing separation.

In a preferred embodiment, in step (5), a portion of the deep hydrogenated diesel is collected as clean diesel.

In a preferred embodiment, in the step (6), the deep hydrogenation diesel oil enters a catalytic cracking unit, is heated (preferably to 180-220 ℃, more preferably 190-210 ℃) and then is sent to a feeding nozzle of a riser of the catalytic cracking unit to be mixed with atomizing steam (preferably, the mass of the atomizing steam is 6-8% relative to the mass of the deep hydrogenation diesel oil); the mixed materials are sprayed out through the feeding nozzle and are contacted with a regenerated catalyst, so that a cracking reaction is carried out to generate light fractions; fractionating the light fraction to obtain catalytic gasoline and catalytic cracking gas, and cutting the catalytic gasoline into catalytic light gasoline and catalytic heavy gasoline; meanwhile, the catalytic cracking gas is subjected to absorption stabilization treatment to recover LPG, and the recovered LPG is separated into mixed carbon three and mixed carbon four. In a further preferred embodiment, in step (6), the regenerated catalyst is a USY type molecular sieve catalyst, such as ZCM-7 molecular sieve catalyst, CHZ-1 molecular sieve catalyst and LCH-7 molecular sieve catalyst; or a composite molecular sieve catalyst, such as RMG catalyst. In a further preferred embodiment, in the step (6), the fractionation is performed under the operating conditions of: the pressure at the top of the fractionating tower is 0.05MpaG to 0.15MpaG, preferably 0.08MPaG to 0.12MpaG, the operation temperature at the top of the fractionating tower is 80 ℃ to 150 ℃, preferably 90 ℃ to 110 ℃, and the operation temperature at the bottom of the fractionating tower is 280 ℃ to 350 ℃, preferably 300 ℃ to 320 ℃. In a further preferred embodiment, in the step (6), the operation conditions of the cutting process are: the temperature of the top of the catalytic gasoline cutting tower is 80-95 ℃, preferably 85-90 ℃, and the pressure of the top of the catalytic gasoline cutting tower is 0.2-0.3 MPaG, preferably 0.22-0.28 MPaG. In a further preferred embodiment, in the step (6), the operation conditions of the absorption stabilization treatment are: the tower top temperature of the catalytic cracking absorption tower is 45-55 ℃, preferably 48-52 ℃, and the tower top pressure is 1.0-1.5 MPaG, preferably 1.2-1.4 MPaG; the tower top temperature of the catalytic cracking desorption tower is 60-70 ℃, preferably 62-68 ℃, and the tower top pressure is 1.0-2.0 MPaG, preferably 1.2-1.4 MPaG; the temperature at the top of the stabilizer column is 60 ℃ to 70 ℃, preferably 62 ℃ to 68 ℃, and the pressure at the top of the stabilizer column is 0.9MPaG to 1.2MPaG, preferably 1MPaG to 1.1 MPaG.

In a preferred embodiment, in the step (6), the catalytic light gasoline enters a light gasoline etherification unit, and after being mixed with methanol, active olefin in the catalytic light gasoline reacts with methanol in the presence of an etherification catalyst to generate an alkyl ether product, and the alkyl ether product is rectified and extracted to separate out methanol, so as to obtain the etherified gasoline. In a further preferred embodiment, in the step (6), the etherification catalyst is a non-basic cation exchange resin catalyst of D001 type, D005 type or D005IIS type. In a further preferred embodiment, in the step (6), the active olefin in the catalytic light gasoline is reacted with methanol under the conditions of 35 ℃ to 55 ℃ (preferably 40 ℃ to 50 ℃), 0.25MPaG to 1.0MPaG (preferably 0.6MPaG to 0.8 MPaG). In a further preferred embodiment, in the step (6), the distillation is performed under the following operating conditions: the temperature of the top of the rectifying tower is 65-80 ℃ (preferably 70-75 ℃), the temperature of the bottom of the rectifying tower is 125-135 ℃ (preferably 128-132 ℃), and the operating pressure of the rectifying tower is 0.2 MPaG-0.3 MPaG (preferably 0.22 MPaG-0.28 MPaG). In a further preferred embodiment, the extraction is carried out under the following operating conditions: the operating temperature of the extraction tower is 35-45 ℃ (38-42 ℃ is preferred), and the operating pressure of the extraction tower is 0.5-1.0 MPaG (0.7-0.9 MPaG is preferred).

In the step (7), the catalytic reforming may be performed using a known continuous catalytic reforming technique or a semi-regenerative catalytic reforming technique. For example, as a preferred embodiment, in the step (7), the catalytic reforming may be performed under the following conditionsThe following were carried out: the mass space velocity is 1.5-2.5 (preferably 1.8-2.2), the temperature is 510-530 ℃ (preferably 515-525 ℃) and the pressure is 0.1-0.5 MPaG (preferably 0.2-0.4 MPaG). In a further preferred embodiment, in the step (7), the catalyst for catalytic reforming is Pt — Al₂O₃。

In the step (7), the naphtha from the tail gas with stable absorption is processed into the reformed gasoline through catalytic reforming, so that the liquid yield can be increased, and the additional value is improved.

In a preferred embodiment, in the step (7), the C5/C6 (carbon five/carbon six) isomerization may be performed using a well-known isomerization technique. As a preferred embodiment, in the step (7), the operation conditions of the isomerization treatment are: at a temperature of 160 ℃ to 190 ℃ (preferably 170 ℃ to 180 ℃) and a pressure of 2.5MPaG to 3.5MPaG (preferably 2.9MPaG to 3.3 MPaG). In a further preferred embodiment, in the step (7), the isomerization-treated catalyst is Pt-Cl/Al₂O₃。

In a preferred embodiment, in the step (7), in order to recover the reformed LPG and the isomerized LPG, absorption stabilizing systems may be provided separately in the catalytic reforming unit or in the C5/C6 isomerization unit, or in a common absorption stabilizing system in either unit (to save equipment investment).

In the step (7), the light naphtha is subjected to C5/C6 isomerization treatment, so that the octane number of the light naphtha (C5/C6 components) can be increased, and the obtained isomerized gasoline can be blended into a high-octane gasoline product as required.

In a preferred embodiment, in step (8), the mixed C.sub.four is mixed with methanol in the MTBE unit, then isobutylene in the mixed C.sub.four reacts with methanol in the presence of a catalyst to produce an MTBE product, and the MTBE product is rectified and extracted to separate out methanol, thereby obtaining MTBE and ether-bound C.sub.four. In a further preferred embodiment, in the step (8), the catalyst is a strong-acid cation exchange resin, preferably a D006 type strong-acid large-pore cation exchange resin. In a further preferred embodiment, in step (8), isobutylene in the mixed C4 is reacted with methanol at 40 ℃ to 70 ℃ (preferably 55 ℃ to 65 ℃), 0.45MPaG to 1.0MPaG (preferably 0.5MPaG to 0.7 MPaG). In a further preferred embodiment, in the step (8), the distillation is performed under the following operating conditions: the temperature of the top of the rectifying tower is 60-80 ℃ (preferably 65-75 ℃), the temperature of the bottom of the rectifying tower is 120-140 ℃ (preferably 125-135 ℃), and the operating pressure of the rectifying tower is 0.55-0.75 MPaG (preferably 0.6-0.7 MPaG). In a further preferred embodiment, in the step (8), the extraction is performed under the following operating conditions: the operating temperature of the extraction tower is 30-50 ℃ (preferably 35-45 ℃), and the operating pressure of the extraction tower is 0.5-0.7 MPaG (preferably 0.55-0.65 MPaG).

In the step (9), the alkylation treatment may be carried out by any known alkylation technique. In a preferred embodiment, in the step (9), the alkylation treatment may be carried out using any one of alkylation techniques selected from the group consisting of: sulfuric acid process alkylation technology, hydrofluoric acid process alkylation technology, solid acid alkylation technology and ionic liquid alkylation technology; preferably, the alkylation treatment may be performed using any one selected from sulfuric acid process alkylation technology, hydrofluoric acid process alkylation technology, and ionic liquid alkylation technology. In a further preferred embodiment, the process conditions of the sulfuric acid process alkylation technique may be: the volume ratio of acid to hydrocarbon (0.5-1.5) is 1, preferably (0.8-1.2) 1, the volume ratio of isobutane to olefin (alkane to olefin ratio) in the reactor feed is 5-15: 1, preferably (8-12) 1, the reaction temperature is 4.5-12 ℃ (preferably 5-10 ℃), the reaction pressure is 0.2-0.8 MPaG (preferably 0.3-0.6 MPaG), and the reaction time is 20-30 min (preferably 22-26 min). In a further preferred embodiment, the process conditions of the hydrofluoric acid process alkylation technique may be: the volume ratio of isobutane to olefin (alkane/alkene ratio) in the reactor feed is (10-15): 1, the reaction temperature is 25-40 ℃, and the reaction time is 5-20 min. In a further preferred embodiment, the process conditions of the ionic liquid alkylation technique may be: the volume ratio of isobutane to olefin (alkane-olefin ratio) in the reactor feed is (5-15): 1, preferably (8-12):1, the reaction temperature is 10-30 ℃ (preferably 15-25 ℃), the reaction pressure is 0.5-1.0 MPaG (preferably 0.6-0.9 MPaG), and the reaction time is 0.3-2.5 min (preferably 1.5-2 min).

In a preferred embodiment, in said step (10), said blending is carried out in a gasoline blending pool. Wherein the blending ratio of the alkylated gasoline, the catalytic heavy gasoline, the etherified gasoline, the reformed gasoline, the isomerized gasoline and the MTBE is not particularly limited, and the ratio can be adjusted according to the production of each unit as required.

In a preferred embodiment, in the step (10), the high octane gasoline is a gasoline having a RON value higher than 93. In still further preferred embodiments, in the step (10), the high octane gasoline is a gasoline that meets the requirements of the national five standards.

The coal pyrolysis product treatment system built according to the technical scheme can produce clean diesel oil, high-octane gasoline products and hydrogen. The gasoline product obtained by the method has the quality not lower than the requirement of the national five standards.

Examples

Materials, reagents and the like used in the following examples are commercially available unless otherwise specified. The experimental procedures used in the following examples are conventional unless otherwise specified.

Example 1

The coal pyrolysis product obtained in this example was obtained by the method described in chinese patent publication No. CN 101717656A. The Shenfu-Dongsheng clean coal 1# is used as a raw material to be made into fuel-oil slurry with solvent oil, the oil slurry is subjected to low and medium pressure hydropyrolysis to obtain a coal pyrolysis product, and the hydropyrolysis process conditions are as follows: pressure 25bar, temperature 400 ℃.

As shown in fig. 1, this example converts the coal pyrolysis product into high-octane gasoline, diesel and hydrogen by the following method, wherein the method comprises the following steps:

(2) the tail gas is firstly washed by a water washing tower to remove ammonia in the tail gas, and then the washed tail gas is cooled to 8 ℃ to separate a small amount of water and heavy hydrocarbon carried in the tail gas, so that heavy hydrocarbon and pretreated tail gas are obtained, and the heavy hydrocarbon is subjected to absorption stabilization treatment to obtain absorption stabilization naphtha and absorption stabilization LPG; the absorption stabilization treatment is carried out by adopting a combination scheme of an absorption desorption tower and a stabilization tower, wherein the theoretical plate number of the absorption desorption tower is 17, the tower top temperature is 45 ℃, the tower bottom temperature is 125 ℃, and the operation pressure is 0.8 MPaG; the number of theoretical plates of the stabilizer is 25, the temperature at the top of the tower is 65 ℃, the temperature at the bottom of the tower is 193 ℃, and the operating pressure is 0.95 MPaG.

(3) The solvent oil enters a solvent oil hydrogenation unit, and the volume space velocity of the solvent oil hydrogenation is 0.5h^-1Under the condition that the volume ratio of the hydrogen oil is 1000, hydrogen supply solvent oil, hydrogenation light oil and hydrogenation gas are obtained after hydrotreating, and the hydrogen supply solvent oil returns to a coal hydrogenation pyrolysis unit;

(4) the bottom-reduced oil residue enters a flexicoking unit, wherein the flexicoking reactor is treated under the conditions that the operating temperature is 510 ℃ and the reactor pressure is 0.1MPaG, and coking heavy oil, coking light oil and heat value are lower than 1200Kcal/Nm³A low heating value gas of (a); mixing the low heating value gas and the pretreated tail gas to perform MDEA desulfurization treatment and pressure swing adsorption treatment to obtain hydrogen;

wherein, the MDEA desulfurization treatment conditions are that the theoretical plate number of the desulfurization tower is 12, the operation pressure is 0.8MPaG, the tower top temperature is 40 ℃, and the amine liquid concentration is 25 wt%; the theoretical plate number of the regeneration tower is 12, the operating temperature is 122 ℃, and the operating pressure is 0.12 MPaG.

Wherein, the processing conditions of the pressure swing adsorption treatment are that the operation temperature of the pressure swing adsorption tower is 40 ℃, the pressure during adsorption is 2.4MPaG, and the pressure during regeneration is 0.045 MpaG; (5) mixing the pyrolysis light oil, the coking light oil, the pyrolysis heavy oil, the coking heavy oil, a part of the absorption-stabilized naphtha and the hydrogenation light oil and the hydrogenation gas, sending the mixture to a deep hydrogenation unit, and carrying out deep hydrogenation on the mixtureThe volume space velocity is 0.5h^-1Under the condition that the volume ratio of hydrogen to oil is 1200, carrying out hydrogenation treatment and then fractionating to obtain deep hydrogenation diesel oil, heavy naphtha, light naphtha and hydrogen-rich tail gas;

(6) collecting a portion of the deeply hydrogenated diesel as clean diesel; the rest part of the deep hydrogenation diesel oil enters a catalytic cracking unit, is heated to 200 ℃ and then is sent to a feeding nozzle of a riser of the catalytic cracking unit to be mixed with atomizing steam, and the mass of the atomizing steam is 8 percent relative to the mass of the deep hydrogenation diesel oil; after mixing, the mixture is sprayed out through a feed nozzle to contact with an LCH-7 type regenerated catalyst from a regenerator in the unit and carry out cracking reaction to generate light fraction; fractionating the light fraction by a fractionating tower in the unit to obtain catalytic gasoline and catalytic cracking gas, and cutting the catalytic gasoline by a cutting tower to obtain catalytic light gasoline and catalytic heavy gasoline; meanwhile, the catalytic cracking gas is absorbed and stabilized by an absorption tower, a desorption tower and a stabilizing tower in sequence to recover the LPG, and the LPG is separated into mixed carbon III and mixed carbon IV;

wherein the fractionation is carried out under the following operating conditions: the pressure at the top of the fractionating tower is 0.1MPaG, the operation temperature at the top of the fractionating tower is 100 ℃, and the operation temperature at the bottom of the fractionating tower is 310 ℃;

the operating conditions of the cutting treatment are as follows: the top temperature of the catalytic gasoline cutting tower is 88 ℃, and the top pressure is 0.25 MPaG; and

the operating conditions of the absorption stabilization treatment are as follows: the tower top temperature of the catalytic cracking absorption tower is 50 ℃, and the tower top pressure is 1.3 MPaG; the temperature at the top of the catalytic cracking desorption tower is 65 ℃, and the pressure at the top of the catalytic cracking desorption tower is 1.5 MPaG; the overhead temperature of the stabilizer is 65 ℃, and the overhead pressure is 1.05 MPaG;

enabling the catalytic light gasoline to enter a light gasoline etherification unit, mixing with methanol, reacting active olefin in the catalytic light gasoline with the methanol at 45 ℃ in the presence of 0.7MPaG and a D005 type non-basic cation exchange resin catalyst to generate an alkyl ether product, and rectifying and extracting the alkyl ether product to separate out the methanol to obtain etherified gasoline;

wherein the rectification operating conditions are as follows: the tower top temperature of the rectifying tower is 73 ℃, the tower bottom temperature is 130 ℃, and the operating pressure of the rectifying tower is 0.25 MPaG; and the operating conditions of the extraction are: the operating temperature of the extraction tower is 40 ℃, and the operating pressure of the extraction tower is 0.8 MPaG;

(7) the heavy naphtha and the remaining portion of the absorbed stable naphtha are fed to a catalytic reforming unit at a mass space velocity of 2.0 and a catalyst of Pt-Al₂O₃And carrying out reforming reactions of aromatization, isomerization and dehydrogenation at a temperature of 520 ℃ and a pressure of 0.3MPaG to obtain reformed gasoline, reformed LPG and reformed gas; the light naphtha is passed to a carbon five/carbon six isomerization unit with a catalyst of Pt-Cl/Al₂O₃And carrying out isomerization reaction at 175 ℃ and 3.1MPaG pressure to obtain isomerized gasoline and isomerized LPG;

in order to recover LPG (reformed LPG and isomerized LPG) produced by the carbon five/carbon six isomerization unit and the catalytic reforming unit, a common absorption stabilizing system is arranged in the catalytic reforming unit;

(8) mixing the mixed C4 with methanol in an MTBE unit, reacting isobutene in the mixed C four with the methanol at 60 ℃ in the presence of a 0.6MPaG and D006 type strongly acidic large-aperture cation exchange resin catalyst to generate an MTBE product, rectifying and extracting the MTBE product to separate out the methanol, and obtaining MTBE and ether C four;

wherein the rectification operating conditions are as follows: the tower top temperature of the rectifying tower is 70 ℃, the tower bottom temperature is 130 ℃, and the operating pressure of the rectifying tower is 0.65 MPaG; and the operating conditions of the extraction are: the operating temperature of the extraction column was 40 ℃ and the operating pressure of the extraction column was 0.6 MPaG;

(9) feeding the etherified C4, the mixed C three, the absorption stable LPG, the reformed LPG and the isomeric LPG into an alkylation unit, and treating by adopting a sulfuric acid method alkylation technology to obtain alkylated gasoline and alkylated LPG;

wherein, the process conditions of the sulfuric acid method alkylation technology are as follows: the volume ratio of acid to hydrocarbon is 1.0:1, the volume ratio of isobutane to olefin (alkane to alkene ratio) in the reactor feed is 10:1, the reaction temperature is 7.3 ℃, the reaction pressure is 0.42MPaG, and the reaction time is 25 min;

(10) blending the alkylated gasoline, the catalytic heavy gasoline, the etherified gasoline, the reformed gasoline, the isomerized gasoline and the MTBE in a gasoline blending pool according to the mass ratio of 0.27:0.1:0.25:0.23:0.1:0.05 to obtain the high-octane gasoline, wherein the RON value, the sulfur content and the nitrogen content of the high-octane gasoline are shown in the following table 1.

Example 2

As shown in fig. 2, the solution of example 1 is generally adopted, with the difference that in step (4), the coking unit adopts the conventional delayed coking technique, which has the process operating conditions: the coke drum overhead pressure was 0.18MPaG and the exit temperature of the feed furnace to the coke drum was 500 ℃. The RON value, sulfur content and nitrogen content of the resulting high octane gasoline are shown in table 1 below.

Example 3

As shown in fig. 3, the technical solution of embodiment 1 is generally adopted, with the difference that: in step (7), the whole of the absorption-stabilized naphtha obtained by off-gas treatment is sent to a catalytic reforming unit (without subjecting part thereof to deep hydrotreating); and in the step (9), the alkylation unit adopts an ionic liquid alkylation technology, and the technology has the process operating conditions that: the volume ratio of isobutane to olefin (alkane to alkene ratio) in the reactor feed was 10:1, the reaction temperature was 20 ℃, the reaction pressure was 0.8MPaG, and the reaction time was 1.8 min. The RON value, sulfur content and nitrogen content of the resulting high octane gasoline are shown in table 1 below.

Table 1 RON number, sulfur content, and nitrogen content of high octane gasoline of examples 1-3

Item	Example 1	Example 2	Example 3
				RON	93.8	93.8	93.7
Sulfur content (mg/kg)	2～5	2～5	2～5
				Nitrogen content (mg/kg)	Is free of	Is free of	Is free of

As can be seen from the above table 1, the RON value of the high-octane gasoline obtained by the method is higher than 93, the sulfur content is 2-5 mg/kg, and the high-octane gasoline does not contain nitrogen and has the quality not lower than the requirements of the national five standards.

Claims

1. A process for converting coal pyrolysis products to high octane gasoline, diesel and hydrogen, wherein the process comprises the steps of:

(8) treating the mixed C4 in an MTBE unit to obtain MTBE and ether C4;

2. The method according to claim 1, wherein in the step (2), the off-gas is first washed by a water washing tower and then cooled to perform the water washing separation.

3. The method of claim 2, wherein in the step (2), the tail gas is firstly washed by a water washing tower and then cooled to 5-15 ℃.

4. The method of claim 3, wherein in the step (2), the tail gas is firstly washed by a water washing tower and then cooled to 6-9 ℃.

5. The method according to any one of claims 1 to 4, wherein in the step (2), the absorption stabilization treatment is carried out using a combination of an absorption desorption column and a stabilization column.

6. The method according to claim 5, wherein in the step (2), the theoretical plate number of the absorption-desorption tower is 10-20, the tower top temperature is 40-50 ℃, the tower bottom temperature is 120-130 ℃, and the operation pressure is 0.5 MpaG-1 MPaG.

7. The process of claim 6, wherein in the step (2), the theoretical plate number of the absorption-desorption tower is 15-18, the tower top temperature is 42-48 ℃, the tower bottom temperature is 124-128 ℃, and the operation pressure is 0.6 MpaG-0.9 MPaG.

8. The method of claim 5, wherein in the step (2), the number of theoretical plates of the stabilizer tower is 20-30, the temperature of the top of the tower is 60-70 ℃, the temperature of the bottom of the tower is 190-200 ℃, and the operating pressure is 0.5 MpaG-1.5 MPaG.

9. The method of claim 8, wherein in the step (2), the number of theoretical plates of the stabilizer column is 22-28, the temperature of the top of the column is 62-68 ℃, the temperature of the bottom of the column is 192-195 ℃, and the operating pressure is 0.9-1.0 MPaG.

10. The process of any one of claims 1 to 4, wherein in step (3), the operating conditions of the solvent oil hydrogenation unit are: solvent oil hydrogenation volume space velocity of 0.1h^-1~1h^-1(ii) a The volume ratio of hydrogen to oil is 500-1500.

11. The method of claim 10, wherein in step (3), the stepThe operating conditions of the solvent oil hydrogenation unit are as follows: solvent oil hydrogenation volume space velocity of 0.5h^-1~1h^-1(ii) a The volume ratio of the hydrogen to the oil is 1000-1500.

12. The process of any one of claims 1-4, wherein in step (3), the hydrogen-donating solvent oil is returned to a coal hydropyrolysis unit.

13. The process of any one of claims 1-4, wherein in step (4), the coking unit is a flexicoking unit.

14. The process of claim 13, wherein in the flexicoking unit, the flexicoking reactor has an operating temperature of from 400 ℃ to 600 ℃; the operating pressure is 0MPaG to 0.2 MPaG.

15. The process of claim 14, wherein the flexicoking reactor is operated at a temperature of 480 ℃ to 550 ℃ in the flexicoking unit; the operation pressure is 0.05MPaG to 0.15 MPaG.

16. The process of claim 13, wherein in step (4), the reduced bottoms sludge is passed to the flexicoking unit and treated to further obtain a heating value of less than 1200Kcal/Nm³Of the low heating value gas.

17. The method according to claim 16, wherein in the step (4), the low heating value gas and the pretreated tail gas are mixed and then subjected to MDEA desulfurization treatment and pressure swing adsorption treatment to obtain hydrogen.

18. The method of claim 17, wherein in the step (4), the conditions of the MDEA desulfurization treatment are as follows: the number of theoretical plates of the desulfurizing tower is 10-20, the operating pressure is 0.5 MpaG-1.0 MPaG, the tower top temperature is 35-45 ℃, and the amine liquid concentration is 20-30 wt%; the theoretical plate number of the regeneration tower is 10-20, the operation temperature is 120-130 ℃, and the operation pressure is 0.10-0.20 MPaG.

19. The method of claim 18, wherein in the step (4), the conditions of the MDEA desulfurization treatment are as follows: the number of theoretical plates of the desulfurizing tower is 10-15, the operating pressure is 0.6 MpaG-0.9 MPaG, the tower top temperature is 38-42 ℃, and the amine liquid concentration is 22-26 wt%; the theoretical plate number of the regeneration tower is 10-15, the operation temperature is 120-125 ℃, and the operation pressure is 0.10-0.15 MPaG.

20. The method according to any one of claims 17 to 19, wherein in the step (4), the pressure swing adsorption process is carried out under the following process conditions: the operation temperature of the pressure swing adsorption tower is less than or equal to 40 ℃, the pressure during adsorption is 2.0 MpaG-2.5 MpaG, and the pressure during regeneration is 0.02 MpaG-0.07 MpaG.

21. The method according to claim 20, wherein in the step (4), the adsorption pressure is 2.3MPaG to 2.5MPaG, and the regeneration pressure is 0.04MpaG to 0.05 MpaG.

22. The process of any one of claims 1-4, wherein in step (4), the coking unit is a delayed coking unit.

23. The process of claim 22, wherein in step (4), the overhead pressure of the coke drum in the delayed coking unit is from 0.15MPaG to 0.2 MPaG; the outlet temperature of the feed heater of the coke drum is from 495 ℃ to 505 ℃.

24. The process of claim 23, wherein in step (4), the overhead pressure of the coke drum in the delayed coking unit is from 0.17MPaG to 0.19 MPaG; the outlet temperature of the feed heater of the coke drum is 498 ℃ to 502 ℃.

25. The method of claim 22, wherein in step (4), the reduced bottoms sludge is passed to the delayed coking unit for further processing to obtain coke.

26. The method of claim 25, wherein the char is gasified and cleaned before being subjected to fischer-tropsch synthesis to produce a fischer-tropsch oil.

27. The process of any one of claims 1-4, wherein in step (4), the coking unit is a fluid coking unit.

28. The method of any one of claims 1-4, wherein in step (5), a portion of the absorption-stabilized naphtha is mixed with the pyrolysis light oil, the coker light oil, the pyrolysis heavy oil, the coker heavy oil, and the hydrogenation light oil and the hydrogenation gas and sent to a deep hydrogenation unit.

29. The process of claim 28, wherein the operating conditions of the deep hydrogenation unit are: the volume space velocity of deep hydrogenation is 0.1h^-1~1h^-1The volume ratio of hydrogen to oil is 500-1500.

30. The process of claim 29, wherein the operating conditions of the deep hydrogenation unit are: the volume space velocity of deep hydrogenation is 0.3h^-1~0.7h^-1The volume ratio of hydrogen to oil is 1000-1400.

31. The method according to any one of claims 1 to 4, wherein in the step (5), the hydrogen-rich tail gas is obtained by fractionation after the hydrotreatment.

32. The method according to claim 31, wherein the hydrogen-rich tail gas is recycled to the step (2) for water washing separation.

33. The process of any one of claims 1-4, wherein in step (5), a portion of the deep hydrogenated diesel is collected as clean diesel.

34. The method according to any one of claims 1 to 4, wherein in the step (6), the deep hydrogenation diesel oil enters a catalytic cracking unit, is heated and then is sent to a feeding nozzle of a riser of the catalytic cracking unit to be mixed with atomizing steam; the mixed materials are sprayed out through the feeding nozzle and are contacted with a regenerated catalyst, so that a cracking reaction is carried out to generate light fractions; fractionating the light fraction to obtain catalytic gasoline and catalytic cracking gas, and cutting the catalytic gasoline into catalytic light gasoline and catalytic heavy gasoline; meanwhile, the catalytic cracking gas is subjected to absorption stabilization treatment to recover LPG, and the recovered LPG is separated into mixed carbon three and mixed carbon four.

35. The process of claim 34 wherein in step (6) said deep hydrogenated diesel is heated in said catalytic cracking unit to from 180 ℃ to 220 ℃.

36. The process of claim 35 wherein in step (6) said deep hydrogenated diesel is heated in said catalytic cracking unit to between 190 ℃ and 210 ℃.

37. The process of claim 34, wherein in step (6), the mass of the atomizing steam is 6% to 8% relative to the mass of the deep hydrogenated diesel oil.

38. The process of claim 34, wherein in step (6), the regenerated catalyst is a USY type molecular sieve catalyst or a composite molecular sieve catalyst.

39. The process of claim 38, wherein the USY-type molecular sieve catalyst is selected from the group consisting of ZCM-7 molecular sieve catalyst, CHZ-1 molecular sieve catalyst and LCH-7 molecular sieve catalyst.

40. The process of claim 38, wherein the composite molecular sieve catalyst is an RMG catalyst.

41. The process of claim 34, wherein in step (6), the fractionation is carried out under the operating conditions of: the tower top pressure of the fractionating tower is 0.05 MpaG-0.15 MpaG; the tower top operating temperature of the fractionating tower is 80-150 ℃; the tower bottom operating temperature of the fractionating tower is 280-350 ℃.

42. The process of claim 41, wherein in step (6), the fractionation is carried out under the operating conditions of: the overhead pressure of the fractionating tower is 0.08 MPaG-0.12 MpaG; the tower top operating temperature of the fractionating tower is 90-110 ℃; the tower bottom operating temperature of the fractionating tower is 300-320 ℃.

43. The method of claim 34, wherein in step (6), the cutting process is performed under the operating conditions of: the tower top temperature of the catalytic gasoline cutting tower is 80-95 ℃; the pressure at the top of the tower is 0.2MPaG to 0.3 MPaG.

44. The method of claim 43, wherein in step (6), the cutting process is performed under the operating conditions of: the tower top temperature of the catalytic gasoline cutting tower is 85-90 ℃; the overhead pressure is 0.22MPaG to 0.28 MPaG.

45. The method according to claim 34, wherein in the step (6), the operation conditions of the absorption stabilization treatment are:

the tower top temperature of the catalytic cracking absorption tower is 45-55 ℃, and the tower top pressure is 1.0-1.5 MPaG;

the tower top temperature of the catalytic cracking desorption tower is 60-70 ℃, and the tower top pressure is 1.0-2.0 MPaG;

the temperature of the top of the stabilizing tower is 60-70 ℃, and the pressure of the top of the stabilizing tower is 0.9-1.2 MPaG.

46. The method according to claim 45, wherein in the step (6), the operation conditions of the absorption stabilization treatment are:

the tower top temperature of the catalytic cracking absorption tower is 48-52 ℃, and the tower top pressure is 1.2-1.4 MPaG;

the tower top temperature of the catalytic cracking desorption tower is 62-68 ℃, and the tower top pressure is 1.2-1.4 MPaG;

the temperature of the top of the stabilizer is 62-68 ℃, and the pressure of the top of the stabilizer is 1-1.1 MPaG.

47. The method of any one of claims 1 to 4, wherein in the step (6), the catalytic light gasoline enters a light gasoline etherification unit, and after the catalytic light gasoline is mixed with methanol, active olefin in the catalytic light gasoline reacts with methanol in the presence of an etherification catalyst to generate an alkyl ether product, and the alkyl ether product is rectified and extracted to separate out methanol, so that the etherified gasoline is obtained.

48. The process of claim 47, wherein in step (6), the etherification catalyst is D001, D005 or D005 type

An S-type non-basic cation exchange resin catalyst.

49. The process of claim 48, wherein in step (6), the etherification catalyst is a D005 type non-basic cation exchange resin catalyst.

50. The method of claim 47, wherein in the step (6), the active olefin in the light catalytic gasoline is reacted with methanol at a temperature of 35 ℃ to 55 ℃ and a pressure of 0.25MPaG to 1.0 MPaG.

51. The method of claim 50, wherein in the step (6), the active olefin in the light catalytic gasoline is reacted with methanol at a temperature of 40 ℃ to 50 ℃ and a pressure of 0.6MPaG to 0.8 MPaG.

52. The method as claimed in claim 47, wherein in the step (6), the operation conditions of the rectification are as follows: the tower top temperature of the rectifying tower is 65-80 ℃; the temperature of the tower bottom is 125-135 ℃; the operating pressure of the rectifying tower is 0.2 MPaG-0.3 MPaG.

53. The method as claimed in claim 52, wherein in the step (6), the operation conditions of the rectification are as follows: the tower top temperature of the rectifying tower is 70-75 ℃; the temperature of the tower bottom is 128-132 ℃; the operating pressure of the rectifying tower is 0.22 MPaG-0.28 MPaG.

54. The method of claim 47, wherein in step (6), the extraction is performed under the following conditions: the operating temperature of the extraction tower is 35-45 ℃; the operating pressure of the extraction tower is 0.5MPaG to 1.0 MPaG.

55. The method of claim 54, wherein in the step (6), the extraction conditions are as follows: the operating temperature of the extraction tower is 38-42 ℃; the operating pressure of the extraction tower is 0.7MPaG to 0.9 MPaG.

56. The process of any one of claims 1 to 4, wherein in step (7), the operating conditions of the catalytic reforming are: the mass space velocity is 1.5-2.5, the temperature is 510-530 ℃, and the pressure is 0.1-0.5 MPaG.

57. The method of claim 56, wherein in step (7), the operating conditions of the catalytic reforming are: the mass airspeed is 1.8-2.2, the temperature is 515-525 ℃, and the pressure is 0.2-0.4 MPaG.

58. The method of claim 56, wherein in step (7), the catalyst for catalytic reforming is Pt-Al₂O₃。

59. The process according to any one of claims 1 to 4, wherein in step (7), the isomerization treatment is carried out under the operating conditions: the temperature is 160-190 ℃; the pressure is 2.5MPaG to 3.5 MPaG.

60. The process of claim 59, wherein in step (7), the isomerization process is carried out under the following conditions: the temperature is 170-180 ℃; the pressure is 2.9MPaG to 3.3 MPaG.

61. The method according to claim 59, wherein in the step (7), the isomerization-treated catalyst is Pt-Cl/Al₂O₃。

62. The process of any one of claims 1 to 4, wherein in step (7), an absorption stabilization system is provided in the catalytic reforming unit or in the C5/C6 isomerization unit, respectively, or in common in either unit.

63. The process of any one of claims 1-4, wherein in step (8), the mixed C4 is mixed with methanol in the MTBE unit, then isobutylene in the mixed C four is reacted with methanol in the presence of a catalyst to produce an MTBE product, and the MTBE product is subjected to rectification and extraction to separate out methanol to obtain MTBE and ether C four.

64. The method of claim 63, wherein in step (8), the catalyst is a strong acid cation exchange resin.

65. The method of claim 64, wherein in step (8), the catalyst is a D006 strongly acidic large pore size cation exchange resin.

66. The process of claim 63, wherein in step (8), the isobutylene in C4 is reacted with methanol at a temperature of 40 ℃ to 70 ℃ and a pressure of 0.45MPaG to 1.0 MPaG.

67. The process of claim 66, wherein in step (8), the isobutylene in C4 is reacted with methanol at a temperature of 55 ℃ to 65 ℃ and a pressure of 0.5MPaG to 0.7 MPaG.

68. The method as claimed in claim 63, wherein in the step (8), the operation conditions of the rectification are as follows: the tower top temperature of the rectifying tower is 60-80 ℃; the temperature of the tower bottom is 120-140 ℃; the operating pressure of the rectifying tower is 0.55MPaG to 0.75 MPaG.

69. The method as claimed in claim 68, wherein in the step (8), the operation conditions of the rectification are as follows: the tower top temperature of the rectifying tower is 65-75 ℃; the temperature of the tower bottom is 125-135 ℃; the operating pressure of the rectifying tower is 0.6 MPaG-0.7 MPaG.

70. The method of claim 63, wherein in the step (8), the extraction is carried out under the following operating conditions: the operating temperature of the extraction tower is 30-50 ℃; the operating pressure of the extraction tower is 0.5MPaG to 0.7 MPaG.

71. The method of claim 70, wherein in the step (8), the extraction is carried out under the following operating conditions: the operating temperature of the extraction tower is 35-45 ℃; the operating pressure of the extraction tower is 0.55MPaG to 0.65 MPaG.

72. The process according to any one of claims 1 to 4, wherein in step (9) the alkylation treatment is carried out using any one of the alkylation techniques selected from: sulfuric acid process alkylation technology, hydrofluoric acid process alkylation technology, solid acid alkylation technology and ionic liquid alkylation technology.

73. The process of claim 72, wherein said alkylation treatment is carried out using any one selected from sulfuric acid process alkylation techniques, hydrofluoric acid process alkylation, ionic liquid alkylation techniques.

74. The process of claim 72, wherein in said step (9), the process conditions of said sulfuric acid process alkylation technique are: the volume ratio of acid to hydrocarbon is (0.5-1.5):1, the volume ratio of isobutane to olefin in the reactor feed is (5-15): 1, the reaction temperature is 4.5-12 ℃, the reaction pressure is 0.2-0.8 MPaG, and the reaction time is 20-30 min.

75. The process of claim 74, wherein in said step (9), the process conditions of said sulfuric acid process alkylation technique are: the volume ratio of acid to hydrocarbon is (0.8-1.2): 1, the volume ratio of isobutane to olefin in the reactor feed is (8-12):1, the reaction temperature is 5-10 ℃, the reaction pressure is 0.3-0.6 MPaG, and the reaction time is 22-26 min.

76. The process of claim 72 wherein in step (9) said hydrofluoric acid process alkylation technique has process conditions: the volume ratio of isobutane to olefin in the reactor feed is (10-15): 1, the reaction temperature is 25-40 ℃, and the reaction time is 5-20 min.

77. The method of claim 72, wherein in step (9), the ionic liquid alkylation technique has process conditions: the volume ratio of isobutane to olefin in the feed of the reactor is (5-15): 1, the reaction temperature is 10-30 ℃, the reaction pressure is 0.5-1.0 MPaG, and the reaction time is 0.3-2.5 min.

78. The method of claim 77, wherein in step (9), the ionic liquid alkylation technique has process conditions of: the volume ratio of isobutane to olefin in the feed of the reactor is (8-12):1, the reaction temperature is 15-25 ℃, the reaction pressure is 0.6-0.9 MPaG, and the reaction time is 1.5-2 min.

79. The method according to any one of claims 1 to 4, wherein in step (10) the blending is carried out in a gasoline blending pool.

80. The method of any one of claims 1-4, wherein the high octane gasoline is a gasoline having a RON value greater than 93.

81. The method of claim 80, wherein said high octane gasoline is a gasoline that meets the requirements of the national five standards.