A method and system for hydrogenating heavy oil using mesoscale reactions

Abstract

The present application relates to the technical field of petroleum chemical industry, disclose a kind of method and system for hydrogenating heavy inferior oil using mesoscale reaction, the method comprises the following steps: (1) after preheating heavy inferior oil, hydrogenolysis is carried out by passing into hydrogen-rich acid gas;Again, pass into catalyst precursor to carry out activation dispersion, obtain slurry;(2) after mixing slurry with saturated hydrogen in feed pipe, from bottom into mesoreactor, mesoreactor from bottom to top is provided with initial hydrogen inlet initial zone, reaction hydrogen inlet reaction zone, delay hydrogen inlet delay zone, carry out hydrogenation reaction, obtain hydrogenation slurry;(3) after slurry separation of hydrogenation slurry, obtain supernatant;(4) after decompression separation of supernatant, diesel and wax oil are obtained respectively.The present application can realize the targeted conversion of supramolecular asphaltene and residual oil, fundamentally slow down the generation and polymerization of coke, solve the problem of long-period operation of full load device.

Description

Technical Field

[0001] This invention relates to the technical field of petrochemicals, and in particular to a method and system for hydrogenating heavy and inferior oils using mesoscale reaction. Background Technology

[0002] Currently, mainstream heavy oil hydrotreating technologies fall into four main categories: fixed bed, moving bed, fluidized bed, and slurry bed (suspension bed). Fixed beds have strict requirements on the density, sulfur content, metal content, and conversion depth of the feedstock. They cannot solve problems such as high pressure drop, short cycle times, and high flux replacement costs caused by asphalt flocculation and sedimentation, and easy coking and blockage of the bed layer. Furthermore, their position is being challenged as the demand for gasoline declines due to the development of the electric vehicle industry. Moving beds have very few industrial applications due to their complex equipment, high investment, and high operational safety risks. Fluidized beds have low total conversion rates, poor asphalt core stability, are prone to coking and blockage, experience frequent downtime due to malfunctions, have high energy consumption, and face difficulties in cleaning and utilizing oil residue. Their market acceptance is declining, and they are gradually being ignored. Typical processes for slurry bed (suspension bed) include VCC, Uniflex, EST, and domestic technologies such as coal + oil co-refining and direct coal liquefaction. For example, CN119286560A discloses a process method and reaction system for residue-oil conversion, which has problems such as reactor bottom vortex overheating, easy separation of product phases, asphalt nucleus flocculation, short operating cycle, high failure rate of key equipment, high energy consumption, high reagent consumption, high methanation, poor raw material adaptability, low yield of target product liquid, and handling of accident safety risks, thus affecting the economy, cleanliness, safety and reliability of the unit.

[0003] The market-oriented slurry-bed (suspension-bed) residue hydrotreating technology comprises two processing units: a slurry reaction unit and a slurry fractionation unit. Industrial-scale reactors have failed to fully resolve the contradictions between thermal cracking and hydrotreating reactions, as well as between monomolecular and bimolecular reactions. The vacuum distillation tower system has not broken free from the constraints of traditional distillation principles, failing to address issues such as mist entrainment and solid-liquid separation. Consequently, the system suffers from multi-point coking difficulties, high hydrogen and catalyst consumption, and low product yield, resulting in weak overall economic performance and dampened investment enthusiasm. In some cases, operating units have had to be modified to add backup vacuum distillation towers for cold switching operation. However, this still suffers from significant drawbacks such as short single-tower operating cycles and a 6-8 day cold switching reaction setback plus 8-10 days for slurry recovery. Ultimately, the nominal design capacity of the unit is only utilized to about 78%. This stopgap measure fails to address the core reaction problems of the key technology and does not completely eliminate the fundamental issues affecting the unit's production cycle and efficient conversion.

[0004] Mesoscale reaction is short for multiphase mesoscale reaction. Mesoscale behavior refers to the complex spatiotemporal structure formed within a system composed of numerous units, ranging from individual units to the global system. In multiphase reactions, this is mainly manifested in the spatiotemporal scale of material structures or interfaces between the molecular scale and particle (including discrete units such as bubbles and droplets, hereinafter the same) scale, and in the spatiotemporal scale of non-uniform structures formed between the particle scale and the reactor scale. The chemical reactions occurring within these reactions (at the atomic and molecular level) exhibit complex behavior due to the influence of mass transfer and diffusion (at the molecular group level) and flow (at the macroscopic statistical level). Summary of the Invention

[0005] To address the aforementioned technical problems, this invention provides a method and system for hydrogenating heavy and inferior oil using mesoscale reaction. It employs a mesoscale reaction design based on mesoscience to create a highly efficient hydrogenation reactor for heavy and inferior oil. A "three-region, two-critical" reaction kinetics and fluid dynamics model is established (gas, liquid, and solid regions; two critical interfaces: solid-liquid and gas-liquid). Based on the phase change regulation of the reaction within the reactor by the mesoscale structure and the efficiency of "three transfers and one reaction (momentum transfer, heat transfer, mass transfer, and chemical reaction)," the synergistic effect of catalyst particle clusters and turbulence within the reactor is studied to improve the rapid targeted reaction of the hydrogenation reactor, increase the lightening rate of heavy and inferior oil and the overflow rate of catalyst and activated hydrogen, reduce the formation of suspended coke and wall coke within the reactor, and slow down the deposition of supramolecular aggregates in the post-slurry fractionation system. Essentially, this solves the problems affecting production stability and economy in slurry bed residue oil hydrogenation units, such as poor feedstock adaptability, low target product yield, high coking rate, and poor anti-interference ability.

[0006] The objective of this invention is achieved through the following technical solution:

[0007] In a first aspect, the present invention provides a method for hydrogenating heavy, low-quality oil using mesoscale reaction, comprising the following steps:

[0008] (1) After preheating the heavy and inferior oil, hydrogen-rich acid gas is introduced to perform hydrogen decomposition; then catalyst precursor is introduced to activate and disperse the oil to obtain a slurry.

[0009] (2) The slurry and saturated hydrogen at 385-395℃ are introduced into the medium reactor from the bottom. The medium reactor is divided into an initial zone for introducing initial hydrogen at 495-505℃, a reaction zone for introducing reactive hydrogen at 495-505℃, and a delay zone for introducing delayed hydrogen at 300-320℃ from bottom to top. The volume flow ratio of saturated hydrogen, initial hydrogen, reactive hydrogen, and delayed hydrogen is 1:4-6:11-13:2-4. The hydrogenation reaction is carried out to obtain hydrogenated slurry.

[0010] (3) After slurry separation, the hydrogenated slurry is used to obtain the supernatant;

[0011] (4) After the supernatant is separated by depressurization, diesel oil and wax oil are obtained respectively.

[0012] After preheating, the heavy, low-quality oil undergoes supramolecular pre-dissociation, completing the molecularization of the clustered asphaltene phase. It then enters the hydrogen dissociation zone and undergoes its first fusion with the microbubbled, hydrogen-rich acidic gas. The hydrogen in the acidic gas reacts with the sulfur in the heavy, low-quality oil to form a fused component of hydrogen sulfide dissolved in the oil. This fused component is then activated and dispersed with the catalyst precursor in the deep activation zone to form a coarse slurry. Simultaneously, hydrogen sulfide reacts with the catalyst precursor to generate a catalyst (molybdenum disulfide), followed by secondary microbubbling. Finally, the highly dispersed phase is split into multiple streams and enters the mixing and slurry-forming zone in a counterclockwise tangential direction, completing high-level hydrogen activation and secondary high-dispersion slurry preparation to obtain a highly dispersed slurry. Heavy and inferior oils are dispersed into slurries through hydrogen decomposition, activation, and mixing, thereby achieving material microfoaming, reducing the activation energy of slurry formation, extending the spatial dimension of slurry preparation, increasing the reaction interface area and net reaction time between hydrogen-rich acidic gas and catalyst precursors, thus improving the degree of molecularization or microfoaming of heavy and inferior oil residues and hydrogen-rich acidic gas, increasing the activation degree of catalyst precursors, and significantly improving the microscopic homogeneity of the slurry.

[0013] The medium reactor is divided into several functional zones—the feed pipe section (before entering the reactor), the initial zone, the reaction zone, and the delay zone—by controlling the temperature and location of the added hydrogen. Hydrogen is the only stream that can provide heat to the reactor during this reaction. Controlling the temperature and location of different hydrogen streams is crucial for controlling the reactor temperature. The separate flow of hydrogen at different temperatures within the reactor allows for precise control of the reaction temperature and the partial pressure of hydrogen in each reaction zone, achieving targeted control of the reaction.

[0014] Saturated hydrogen mixes with the slurry at the feed pipe section, ensuring the slurry entering the reactor is hydrogen saturated. The feedstock oil in the slurry reacts immediately upon entering the reactor, shortening the hydrogen diffusion time and fully utilizing the bottom space of the reactor. Initial hydrogen provides the initial heat for the reaction in the bottom stream and fills the bottom of the reactor with hydrogen, avoiding the risk of coking due to low hydrogen content in the lower part of the reactor and preventing solid precipitation due to insufficient disturbance. Reaction hydrogen mainly provides heat for the hydrogenation reaction, ensuring the main stream entering the reactor reaches the reaction temperature, while continuing to increase the disturbance of the mixture flow and providing kinetic energy from bottom to top for the total stream. Delayed hydrogen is hydrogen below the reaction temperature, thereby reducing the temperature rise caused by exothermic reactions in the early stages, reducing coking due to high temperatures, and increasing the partial pressure of dissolved hydrogen and the steady-state flow around the reactor. The different locations for hydrogen introduction at different temperatures are set according to their different functions, reducing the risk of local overheating in the reactor, achieving targeted conversion of supramolecular asphaltene and residual oil (products with high oil content, high light content, and low methane), and reducing the amount of suspended coke and wall phase coke generated in the reactor, thereby fundamentally reducing the generation and polymerization of subsequent coke.

[0015] The flow rate of saturated hydrogen can saturate the feedstock oil in the slurry. Too much hydrogen will result in a two-phase flow of gas and liquid, while too little hydrogen will not achieve saturation and will need to be replenished later in the reactor, wasting bottom space and reaction time. The initial hydrogen flow rate can provide the initial heat for the reaction and fill the bottom of the reactor with hydrogen. Too much hydrogen will interfere with the flow pattern of the reacting hydrogen and cause the bottom temperature of the reacting hydrogen to be too high. Too little hydrogen will cause coking due to insufficient hydrogen at the bottom of the reactor, and solid phase will deposit at the bottom of the reactor. The flow rate of reacting hydrogen can ensure that the temperature of the main reactant stream reaches the efficient reaction temperature range and that the hydrogen partial pressure is within the efficient reaction range. If the hydrogen flow rate is low, less heat will be carried into the reactor, and the reaction temperature will not reach the efficient reaction range. If the hydrogen flow rate is too high, the reactor temperature will be too high, the reaction will be more aggressive, and energy will be wasted. The flow rate of delayed hydrogen can control the reaction temperature in the delayed zone at the upper edge of the efficient reaction temperature range. If the flow rate is low, the temperature control effect will not be as expected. If the flow rate is too high, the reaction will be dragged into the low-temperature reaction zone, reducing the reaction efficiency.

[0016] Preferably, in step (1), the preheating temperature is 350-360℃.

[0017] The preheating temperature is the highest temperature range before the reaction occurs. If the temperature exceeds this range, the material is more likely to undergo polymerization. If the temperature is below this range, more heat needs to be introduced by hydrogen to continue heating, which increases the amount of hydrogen circulating in the reaction system and increases energy consumption.

[0018] Preferably, in step (1), the hydrogen-rich acidic gas includes hydrogen sulfide, hydrogen, and organic gases, wherein the volume percentage of hydrogen is not less than 80%, and the organic gases include methane and / or ethane; the volume flow ratio of the hydrogen-rich acidic gas to the heavy inferior oil is 100:650-700, and the volume of hydrogen sulfide accounts for 3000-5000 ppm of the heavy inferior oil.

[0019] Preferably, in step (1), the catalyst precursor is molybdenum isooctanoate; the amount of catalyst precursor introduced is such that the mass ratio of catalyst (molybdenum disulfide) in the slurry is 2100-4500 ppm, more preferably 2300-2500 ppm.

[0020] Preferably, in step (1), when the entire system is circulated, the circulating material and the catalyst precursor are introduced together for activation and dispersion. The circulating material includes the catalyst and polycyclic aromatic hydrocarbons. The amount of circulating material and catalyst precursor introduced is such that the mass ratio of catalyst (molybdenum disulfide) in the slurry is 2100-4500 ppm, more preferably 2300-2500 ppm.

[0021] Before entering the reactor, the maximum concentration of the catalyst (molybdenum disulfide) in the slurry is 4500 ppm. Increasing the concentration further would put additional pressure on the system equipment and safety, and since molybdenum disulfide is a solid phase, excessively high concentrations would significantly increase the likelihood of coking. Conversely, if the catalyst concentration in the slurry is too low, below 2100 ppm, insufficient catalyst content in the reactor would result in less hydrogenation and increased cracking reactions, posing a risk of rapid temperature rise and fast coking. A more preferable concentration is 2300-2500 ppm, within which the possibility of coking in the reactor can be minimized.

[0022] Preferably, in step (2), the volume flow ratio of the slurry to saturated hydrogen is 15:10-12; the initial zone, reaction zone, and delay zone account for 10-15%, 50-60%, and 30-35% of the volume of the reactor, respectively; the total time of the hydrogenation reaction is 300-360 min, and the residence time in the initial zone, reaction zone, and delay zone accounts for 10-15%, 50-60%, and 30-35%, respectively.

[0023] Preferably, in step (3), the slurry separation includes high-pressure momentum separation and low-pressure momentum separation performed sequentially.

[0024] Preferably, the high-pressure momentum separation includes: passing the hydrogenated slurry and quenched oil together into a high-pressure momentum separator, and then passing in disturbing hydrogen at the bottom of the separator for separation to obtain a hydrogen-rich oil-gas mixture and a solid-liquid phase slurry.

[0025] Preferably, the temperature of the quench oil is 175-190℃; the volumetric flow rate ratio of the hydrogenation slurry to the quench oil is 10-12:1; the height-to-diameter ratio of the high-pressure momentum separator is 3-4:1, and the feed level is 40-60%; the separation pressure is 15.2-15.75 MPa, and the temperature is 410-415℃; the temperature of the turbulent hydrogen is 380-400℃; and the volumetric flow rate ratio of the turbulent hydrogen to the hydrogenation slurry is 5:3-4.

[0026] High-pressure momentum separation involves the solid-liquid-gas three-phase separation of high-temperature reaction effluents. Therefore, coking is prone to occur under high pressure and high temperature conditions. To reduce the formation of wall-phase coke in the feed line and suspended coke / wall-phase coke in the high-pressure momentum separator, momentum is further increased. Optimizing the height-to-diameter ratio and feed level of the high-pressure momentum separator controls the critical settling velocity of the mist, while simultaneously improving the stability and separation efficiency of the high-pressure momentum separator. The temperature of the hydrogenation slurry is generally 420-430℃. By introducing quenching oil into the feed line of the high-pressure momentum separator, most of the material in the feed line is converted in the low-temperature zone (390-420℃), and a small portion is converted in the delayed-time zone (420-430℃). This controls the temperature of the material in the pipeline, segmenting the coking reaction and reducing coking. Simultaneously, lowering the feed line temperature increases the liquid phase content of the inlet material, altering the phase composition ratio of the feed to the high-pressure momentum separator, further improving the separation efficiency. Disturbance-inducing hydrogen is introduced into the lower part of the high-pressure momentum separator. The hydrogen gas can generate disturbance in the lower part of the high-pressure momentum separator, blowing up the solid phase attached to the tube wall and reducing coking.

[0027] Preferably, in step (3), the low-pressure momentum separation includes: passing the solid-liquid phase mixed slurry into a low-pressure momentum separator, and obtaining heavy components after separation; and obtaining supernatant by flash evaporation and cyclone separation of the heavy components.

[0028] Preferably, the feed level in the low-pressure momentum separator is 40-60%; the separation pressure is 0.4-0.5 MPa, and the temperature is 370-380℃.

[0029] Low-pressure momentum separation is a secondary separation of the bottom solid-liquid phase separated by high-pressure momentum separation. It mainly involves reducing the separation pressure and temperature, and taking advantage of the difference in boiling point and volatility between the slurry and the product oil to separate the light components in advance. The top of the separator contains the light hydrocarbon components, and the bottom of the separator contains the heavy components (heavy reaction products and dispersed solid phase).

[0030] Preferably, in step (3), the temperature of the flash evaporation is 370-380℃ and the pressure is 0.2-0.25MPa; the temperature of the cyclone separation is 365-375℃ and the pressure is 1.1-1.2MPa.

[0031] Flash evaporation is a secondary separation of light hydrocarbons from the heavy components. Flash evaporation is carried out in a flash evaporator, and by using a hot-cut flash evaporator, the aggregation and formation of suspended coke and wall coke are reduced. To reduce the coking rate in the solid-liquid separation vacuum tower, a cyclone separator is installed between the flash evaporator and the solid-liquid separation vacuum tower. Utilizing the feed's swirling flow and density difference from the bottom oil of the flash evaporator, it separates the oil into a bottom heavy slurry and a top supernatant. The heavy slurry is returned to the deep activation zone for recycling, while the supernatant is sent to the solid-liquid separation vacuum tower for product separation. Pre-separation of heavy components in the slurry reduces the feed load to the solid-liquid separation vacuum tower, decreases the amount of heavy components in the feed, and significantly alleviates the pressure on the solid-liquid separation vacuum tower.

[0032] Preferably, in step (4), the solid-liquid separation is carried out in a solid-liquid separation pressure reducing tower. The solid-liquid separation tower consists of an upper narrow tower section, a middle wide tower section, and a lower conical section from top to bottom. The residence times in the middle and lower sections are 270-300s and 60-90s, respectively, and the total residence time in the middle and lower sections is 5.5-6.5min. The tower top temperature is 52-62℃, the tower bottom temperature is 305-315℃, and the tower pressure is -0.075MPa to -0.065MPa. Two side streams are used for extraction, with diesel products extracted from the upper section and wax oil products extracted from the middle section.

[0033] The solid-liquid separation vacuum tower features a design with packing at the top, spraying in the middle, and a conical section at the bottom. It incorporates a multi-section variable-diameter design (thicker in the middle and thinner at both ends) to reduce the probability of suspended coke and wall coke accumulation through adhesion between the vacuum tower and pipelines. Separation begins upon material entry into the vacuum tower. By adjusting the residence time of the material in different sections, the critical settling velocity of solids is altered, reducing solid entrainment in the liquid phase. The solid and heavy liquid phases fall to the bottom of the vacuum tower and are recycled to the deep activation zone for reuse. This avoids solid entrainment in the liquid phase clogging the packing at the top of the tower, significantly reducing the rate of packing blockage. Simultaneously, reducing the contact interface between the material in the middle section and the tower internals and pipe walls, and increasing the flow velocity of the material in the bottom conical section, reduces the likelihood of coke accumulation and coke blockage in the lower and middle sections of the vacuum tower.

[0034] Secondly, the present invention provides a system for mesoscale reactive hydrogenation of heavy and inferior oil using the above-mentioned method, comprising a feedstock initial reaction device, a meso reactor, a slurry separation device, and a vacuum separation device connected in sequence; the feedstock initial reaction device includes a de-disintegration preheater, a hydrogen de-disintegration reactor, and a deep activation reactor connected in sequence; the meso reactor has a slurry inlet at the bottom and a slurry outlet at the top; the slurry inlet is connected to the deep activation reactor through a feed pipe, and the feed pipe is also provided with a saturated hydrogen inlet; the meso reactor includes an initial zone, a reaction zone, and a delay zone arranged from bottom to top, the initial zone has an initial hydrogen inlet, the reaction zone has a reaction hydrogen inlet, and the delay zone has a delay hydrogen inlet.

[0035] Preferably, the feedstock initial reaction unit includes a desiccant preheater, a hydrogen desiccant reactor, a deep activation reactor, and a disperser connected in sequence; the desiccant preheater is equipped with a heavy and inferior oil inlet; the hydrogen desiccant reactor is equipped with a hydrogen-rich acid gas inlet; the deep activation reactor is equipped with a catalyst precursor inlet and a circulation pipeline inlet; the outlet of the disperser is connected to the slurry inlet of the medium reactor.

[0036] Preferably, the slurry separation device includes a high-pressure slurry device and a low-pressure slurry device.

[0037] Preferably, the high-pressure slurry unit includes a high-pressure momentum separator and a cooling device; the lower part of the high-pressure momentum separator is provided with a disturbance-inducing hydrogen inlet, and the middle part is provided with a slurry inlet, which is connected to the slurry outlet of the medium reactor and the quench oil pipeline respectively; the bottom of the high-pressure momentum separator is provided with a solid-liquid phase outlet, which is connected to the cooling device and the low-pressure slurry unit in sequence; the top of the high-pressure momentum separator is provided with a hydrogen-rich oil and gas outlet.

[0038] Preferably, the low-pressure slurry device includes a low-pressure momentum separator, a flash tower and a cyclone separator connected in sequence, and also includes an oil washing tower connected to the low-pressure momentum separator and the flash tower respectively; the bottom of the cyclone separator is provided with a heavy slurry outlet, which is connected to a circulation pipeline.

[0039] Preferably, the solid-liquid vacuum separation device includes a solid-liquid vacuum separation tower; the solid-liquid vacuum separation tower consists of an upper narrow tower section, a middle wide tower section, and a lower conical section from top to bottom; the middle wide tower section is connected to a cyclone separator and an oil washing tower, and a wax oil product outlet is provided on the side line of the middle wide tower section; a diesel oil outlet is provided on the side line of the upper narrow tower section; and a heavy wax oil outlet is also provided at the bottom of the solid-liquid vacuum separation tower, which is connected to a circulation pipeline.

[0040] Dual-circulation slurry circulation mainly refers to the merging and recycling of heavy slurry separated by the hydrocyclone separator and heavy wax oil separated from the bottom of the solid-liquid vacuum separation tower into the deep activation reactor of the feedstock initial reaction system for reuse. The heavy residue oil at the bottom of the solid-liquid vacuum separation tower, due to its high catalyst metal content, maintains a constant solids content in the vacuum tower circulating oil by controlling the discharge rate of the residue oil.

[0041] Compared with the prior art, the present invention has the following beneficial effects:

[0042] (1) Through the precise temperature control and time delay setting of the medium-scale reactor, the reaction selectivity and cracking conversion efficiency are improved, and the targeted conversion of supramolecular asphaltene and residue oil is realized. The reaction has high oil production rate, high light content and low methane content.

[0043] (2) By using a mesoscale reactor and high-pressure circulating wash oil to stop the reaction process, the original conditions for coke formation are blocked or suppressed, thus fundamentally slowing down the formation and polymerization of coke.

[0044] (3) By setting up a hot-cut double slurry balance flash tower, a cyclone separator and a multi-stage variable diameter solid-liquid separation decompression tower, the concentration of small coke lumps is reduced, the environment for the adhesion and enrichment of coke lumps is avoided, the possibility of coke lumps being deposited and enriched is reduced, and the problem of full-load and long-cycle operation of the system is finally solved. Attached Figure Description

[0045] Figure 1 This is a schematic diagram of the structure of the feedstock oil initial reaction device, the medium reactor, and the high-pressure slurry device in the system of the present invention.

[0046] Figure 2 This is a schematic diagram of the low-pressure slurry device and the pressure-reducing separation device in the system of the present invention.

[0047] The attached diagram is labeled as follows: 1. De-disintegration preheater; 101. Heavy and inferior oil pipeline; 2. Hydrogen de-disintegration reactor; 201. Hydrogen-rich acidic gas pipeline; 3. Deep activation reactor; 301. Catalyst precursor pipeline; 4. Disperser; 5. Medium reactor; 501. Medium reactor feed pipeline; 502. Saturated hydrogen pipeline; 503. Initial hydrogen pipeline; 504. Reactive hydrogen pipeline; 505. Delayed hydrogen pipeline; 6. High-pressure momentum separator; 601. 602. Quenching oil pipeline; 603. Separator feed pipeline; 604. Hydrogen-enhancing pipeline; 605. Hydrogen-rich oil and gas; 606. Solid-liquid mixed slurry; 7. Low-pressure momentum separator; 8. Flash distillation tower; 9. Cyclone separator; 10. Oil washing tower; 11. Solid-liquid vacuum separation tower; 1101. Upper narrow tower section; 1102. Middle wide tower section; 1103. Lower conical section; 12. Circulation pipeline; 13. Diesel products; 14. Wax oil products. Detailed Implementation

[0048] The technical solution of the present invention will be illustrated below with specific embodiments, but the scope of protection of the present invention is not limited thereto.

[0049] 1. System

[0050] The system for hydrogenating heavy and inferior oil using mesoscale reaction includes a feedstock initial reaction unit, a mesoreactor 5, a slurry separation unit, and a vacuum separation unit connected in sequence.

[0051] 1.1 Initial reaction unit for feedstock oil

[0052] The feedstock initial reaction unit includes a de-disintegration preheater 1, a hydrogen de-disintegration reactor 2, a deep activation reactor 3, and a disperser 4 connected in sequence. The de-disintegration preheater 1 has a heavy and inferior oil inlet connected to the heavy and inferior oil pipeline 101, and its outlet is connected to the hydrogen de-disintegration reactor 2. The hydrogen de-disintegration reactor 2 is equipped with a gas-liquid mixer and a hydrogen-rich acid gas inlet connected to the hydrogen-rich acid gas pipeline 201. The introduced hydrogen-rich acid gas mixes and reacts with the heavy and inferior oil through the gas-liquid mixer. The outlet of the hydrogen de-disintegration reactor 2 is connected to the deep activation reactor 3. The deep activation reactor 3 has a catalyst precursor inlet and a circulation pipeline inlet. The catalyst precursor inlet is connected to the catalyst precursor pipeline 301 to introduce the catalyst precursor (molybdenum isooctanoate), and the circulation pipeline inlet is connected to the circulation pipeline 12 to introduce the circulating slurry (including continuous phase polycyclic aromatic hydrocarbons and catalyst). The deep activation reactor 3 is equipped with a multiphase mixer. The catalyst precursor and circulating slurry are mixed and reacted with heavy, low-quality oil containing hydrogen-rich acidic gas through the multiphase mixer. The outlet of the deep activation reactor 3 is connected to the disperser 4. The disperser 4 is equipped with a swirl flow component. The slurry enters the disperser 4 in multiple counterclockwise tangential streams, forming a slurry through swirl flow. Furthermore, a slurry shear transfer pump is installed on the pipeline connecting the outlet of the disperser 4 to the slurry inlet of the intermediate reactor 5. The slurry is also returned to the middle of the disperser 4 through the slurry shear transfer pump, forming a slurry through swirl flow again, thereby improving the dispersion of the slurry.

[0053] 1.2 Mesoreactor

[0054] The intermediate reactor 5 has a slurry inlet at the bottom and a slurry outlet at the top. The slurry inlet is connected to the outlet of the disperser 4 through the intermediate reactor feed pipe 501. The intermediate reactor feed pipe 501 has a saturated hydrogen inlet, which is connected to a saturated hydrogen pipe 502 to introduce saturated hydrogen. The saturated hydrogen and slurry are initially mixed in the intermediate reactor feed pipe 501. The intermediate reactor 5 includes an initial zone, a reaction zone, and a delay zone arranged sequentially from bottom to top. The initial zone has an initial hydrogen inlet connected to the initial hydrogen pipe 503, the reaction zone has a reaction hydrogen inlet connected to the reaction hydrogen pipe 504, and the delay zone has a delay hydrogen inlet connected to the delay hydrogen pipe 505.

[0055] 1.3 Slurry Separation Unit

[0056] The slurry separation device includes a high-pressure slurry device and a low-pressure slurry device.

[0057] The high-pressure slurry unit includes a high-pressure momentum separator 6 and a cooling device. The high-pressure momentum separator 6 has a slurry inlet in its middle section, which is connected to the slurry outlet of the intermediate reactor 5 via a separator feed pipe 601. A quench oil inlet is located on the separator feed pipe 601, connected to a quench oil pipe 602 to introduce quench oil. The quench oil mixes with the hydrogenated slurry in the separator feed pipe 601 to reduce the temperature of the material entering the high-pressure momentum separator 6. A disturbance-inducing hydrogen inlet is located at the bottom of the high-pressure momentum separator 6, connected to a disturbance-inducing hydrogen pipe 603. By introducing hydrogen gas, disturbance is generated, blowing away the solid phase adhering to the pipe wall and reducing coking. A solid-liquid phase outlet is located at the bottom of the high-pressure momentum separator 6, connected to the cooling device. The solid-liquid phase mixed slurry 605 exiting the solid-liquid phase outlet is cooled before being introduced into the low-pressure slurry unit. The top of the high-pressure momentum separator 6 is equipped with a hydrogen-rich oil and gas outlet. The hydrogen-rich oil and gas 604 is discharged from the hydrogen-rich oil and gas outlet. After the hydrogen-rich oil and gas 604 obtained after high-pressure separation is first removed from light hydrocarbons and cooled for desulfurization, the hydrogen is pressurized by the compressor to supply hydrogen to the reaction system.

[0058] The low-pressure slurry unit includes a low-pressure momentum separator 7, a flash tower 8, a cyclone separator 9, and an oil washing tower 10. The solid-liquid phase outlet of the high-pressure momentum separator 6 is connected to the inlet of the low-pressure momentum separator 7. The solid-liquid mixed slurry 605 is cooled before being introduced into the low-pressure momentum separator 7. The bottom outlet of the low-pressure momentum separator 7 is connected to the flash tower 8. The flash tower 8 is a dual-tower hot-cut structure, which reduces the aggregation and formation of suspended coke and wall phase coke. Specifically, two identical flash towers 8 are arranged in parallel. The flash tower 8 is equipped with a preheating process, warm pipeline, and isolation blind plate to ensure that the two towers can switch heat online without disturbance. The flash tower 8 reduces the contact interface between the slurry and the tower internals by increasing the spacing between the upper trays, reducing the number of lower trays, and enlarging the aperture of the internal components of the trays. This reduces the adhesion area of ​​the coke or the increased adhesion area of ​​the polymer, ultimately reducing the formation of coke lumps, especially large coke lumps. When the fluctuation range of the outlet flow or pressure of the operating flash tower 8 increases, preparations should be made to switch flash tower 8.

[0059] The undisturbed heat switching steps of flash tower 8 are as follows:

[0060] A. The oil from the outlet of the fourth reducing pump is introduced into the standby flash tower for flushing and heating to establish the liquid level in the standby flash tower. When the liquid level is high, the material is returned to the inlet of the fourth reducing pump, and the standby flash tower is preheated to 150°C.

[0061] B. The operating tower oil is introduced into the standby flash tower through the heating tower line to continue heating. When the temperature of the standby flash tower reaches 280°C, the main pipeline material is gradually introduced into the standby flash tower to reach 350°C.

[0062] C. Adjust the feed and temperature of the standby flash tower. When the operating parameters of the standby flash tower are close to the parameters of the original operating flash tower, both towers will run simultaneously.

[0063] D. Gradually switch the original operating flash tower to the standby flash tower, and close the manual valve from the original operating flash tower to the cyclone separator;

[0064] E. After shutting down the original flash evaporator, perform oil removal and water washing replacement in preparation for coke removal and impurity removal.

[0065] The outlet at the bottom of the flash distillation tower 8 is connected to the hydrocyclone 9. After hydrocyclone separation, the supernatant from the top outlet of the hydrocyclone 9 is fed into the solid-liquid vacuum separation tower 11. The bottom of the hydrocyclone 9 has a heavy slurry outlet, which is connected to the circulation pipe 12. The oil washing tower 10 is connected to the top outlet of the low-pressure momentum separator 7 and the top outlet of the flash distillation tower 8. The oil washing tower 10 mainly separates the gasoline component and lighter hydrocarbons from the diesel component in the light hydrocarbons. After separation, the product is fed into the solid-liquid vacuum separation tower 11 from the bottom outlet of the oil washing tower 10, thereby improving the product yield.

[0066] 1.4 Solid-Liquid Vacuum Separation Device

[0067] The solid-liquid vacuum separation device includes a solid-liquid vacuum separation tower 11, a vacuum pumping device, and a heat tracing device. The top of the solid-liquid vacuum separation tower 11 is connected to the vacuum pumping device. The solid-liquid vacuum separation tower 11 consists of an upper narrow tower section 1101, a middle wide tower section 1102, and a lower conical section 1103 from top to bottom. The diameters of the upper narrow tower section 1101, the middle wide tower section 1102, and the lower conical section 1103 are 4-9m, 8-12m, and 4-8m, respectively. The internal components of the upper narrow tower section 1101 are macroporous packing, the internal components of the middle wide tower section 1102 are empty tower spray trays, and the internal components of the lower conical section 1103 are empty tower trays. The lower conical section 1103 is equipped with a low-pressure steam inlet and a circulating quench slurry inlet. The liquid phase section in the lower conical section 1103 is also connected to the heat tracing equipment. By maintaining the temperature of the entire heat tracing zone, the viscosity of the liquid phase section at the bottom of the tower and the pipeline is controlled, preventing the laminar flow effect on the pipe wall, reducing the agglomeration and enrichment of coke lumps, and reducing the occurrence of coke extending from the pipe wall to the center of the pipeline due to the laminar flow effect in the bottom conical section and the discharge pipeline. This ensures the full-load long-term operation of the solid-liquid separation pressure reducing tower 11.

[0068] The middle wide section 1102 is connected to the cyclone separator 7 and the oil washing tower 10, and a wax oil product outlet is provided on the side of the middle wide section 1102 for collecting wax oil product 14. A diesel product outlet is provided on the side of the upper narrow section 1101 for collecting diesel product 13. The bottom of the lower conical section 1103 of the solid-liquid vacuum separation tower 11 is also provided with a heavy wax oil outlet, which is connected to the circulation pipeline 12.

[0069] 2. Methods

[0070] The method for hydrogenating heavy, low-quality oil using the above system includes the following steps:

[0071] (1) After preheating the heavy and inferior oil in the de-disintegration preheater, the temperature of the heavy and inferior oil is 350-360℃; the heavy and inferior oil is introduced into the hydrogen de-disintegration reactor, and hydrogen-rich acid gas (including hydrogen sulfide and hydrogen) is introduced at the same time for hydrogen de-disintegration. The volume flow ratio of hydrogen-rich acid gas to heavy and inferior oil is 100:650-700; then it is introduced into the deep activation reactor, and a catalyst precursor (molybdenum isooctanoate) or a catalyst precursor (molybdenum isooctanoate) and circulating slurry (including polycyclic aromatic hydrocarbons and catalyst) are introduced for deep activation. The amount introduced is such that the mass ratio of catalyst (molybdenum disulfide) in the slurry is 2100-4500ppm; the material after deep activation is introduced into the disperser and dispersed into a slurry to obtain a highly dispersed slurry.

[0072] (2) The slurry is mixed with saturated hydrogen at 385-395℃ in the feed pipe at a volume flow ratio of 15:10-12. The mixture is then introduced into the intermediate reactor through the bottom inlet. The intermediate reactor is divided into an initial zone for introducing initial hydrogen at 495-505℃, a reaction zone for introducing reactive hydrogen at 495-505℃, and a delay zone for introducing delayed hydrogen at 300-320℃. The initial zone, reaction zone, and delay zone occupy the following volume percentages of the intermediate reactor: The volumetric flow rates of saturated hydrogen, initial hydrogen, reactive hydrogen, and delayed hydrogen are 10-15%, 50-60%, and 30-35%, respectively, with a ratio of 1:4-6:11-13:2-4. The hydrogenation reaction is carried out in a medium reactor for a total time of 300-360 min. The residence time of the material in the initial zone, reaction zone, and delayed zone is 10-15%, 50-60%, and 30-35%, respectively. The hydrogenated slurry is then obtained from the top outlet of the medium reactor.

[0073] (3) The hydrogenated slurry and quench oil at 175-190℃ are mixed in the feed pipe. The volume flow ratio of the hydrogenated slurry to the quench oil is 10-12:1. The mixture is then introduced into a high-pressure momentum separator at the middle inlet of the separator. The height-to-diameter ratio of the high-pressure momentum separator is 3-4:1, and the feed liquid level is 40-60%. At the same time, hydrogen at 380-400℃ is introduced into the lower part of the high-pressure momentum separator. The volume flow ratio of the hydrogen to the hydrogenated slurry is 5:3-4. The separation is carried out at a pressure of 15.2-15.75MPa and a temperature of 410-415℃ to obtain a hydrogen-rich oil-gas and solid-liquid mixed slurry. After cooling, the solid-liquid mixture is fed into a low-pressure momentum separator at a liquid level of 40-60% for separation at a pressure of 0.4-0.5 MPa and a temperature of 370-380℃. Heavy components are obtained at the bottom of the low-pressure momentum separator. These heavy components are then fed into a flash evaporator at a temperature of 370-380℃ and a pressure of 0.2-0.25 MPa. The material exiting from the bottom outlet of the flash evaporator enters a hydrocyclone separator for further separation at a temperature of 365-375℃ and a pressure of 1.1-1.2 MPa. Supernatant is obtained at the top of the hydrocyclone separator, and heavy slurry is obtained at the bottom. This heavy slurry is used as a circulating slurry and fed into a deep activation reactor for recycling.

[0074] (4) The supernatant is fed into the middle section of the solid-liquid separation vacuum tower. At the same time, the oil washing tower separates the light components in the low-pressure inertial momentum separator and the flash tower. The temperature of the oil washing tower is 260-270℃ and the pressure is 0.15-0.18MPa. The heavy components obtained after separation are fed into the middle section of the solid-liquid vacuum separation tower from the bottom outlet of the oil washing tower. Vacuum separation is carried out in the solid-liquid vacuum separation tower. The temperature at the top of the tower is 52-62℃, the temperature at the bottom of the tower is 305-315℃, and the pressure is -0.075MPa to -0.065MPa. The residence time in the middle section and the lower section is 270-300s and 60-90s, respectively. The total residence time in the middle section and the lower section is 5.5-6.5min. Two side streams are collected. Diesel products are collected from the upper section and wax oil products are collected from the middle section. In addition, heavy wax oil is obtained at the bottom of the solid-liquid vacuum separation tower. The heavy wax oil is used as a circulating slurry and is fed into the deep activation reactor for recycling.

[0075] Example 1

[0076] Heavy and inferior oils: These include mixtures of vacuum residue, heavy aromatics, and hydrotreated tail oil, characterized by high specific gravity (1.05), high viscosity (7500 cSt (100℃)), high residual carbon content (26 wt%), high sulfur content (above 6.5 wt%), high nitrogen content (above 5000 ppm), high metal content (Ni>90 ppm, V>260 ppm), and high asphaltene content (19.5 wt%).

[0077] The method for hydrogenating heavy, low-quality oil using the above system includes the following steps:

[0078] (1) After preheating the heavy and inferior oil in the de-disintegration preheater, the temperature of the heavy and inferior oil is 356℃; the heavy and inferior oil is fed into the hydrogen de-disintegration reactor, and hydrogen-rich acid gas (including hydrogen sulfide and hydrogen) is fed into the reactor for hydrogen de-disintegration. The volume flow ratio of hydrogen-rich acid gas to heavy and inferior oil is 100:675; then it is fed into the deep activation reactor, and catalyst precursor (molybdenum isooctanoate) and circulating slurry (including polycyclic aromatic hydrocarbons and catalyst) are fed into the reactor for deep activation. The mass ratio of catalyst precursor to circulating slurry is 1:700, and the mass ratio of catalyst (molybdenum disulfide) in the slurry is 2300-2500ppm; the material after deep activation is fed into the disperser and dispersed into a slurry to obtain a highly dispersed slurry.

[0079] (2) The slurry and 391℃ saturated hydrogen are mixed in the feed pipe. The volume flow ratio of the slurry to the saturated hydrogen is 15:11. The mixture is then introduced into the medium reactor through the bottom inlet of the reactor. The medium reactor is divided into an initial zone for introducing 501℃ initial hydrogen, a reaction zone for introducing 501℃ reaction hydrogen, and a delay zone for introducing 312℃ delayed hydrogen from bottom to top. The initial zone, reaction zone, and delay zone account for 10%, 55%, and 35% of the volume of the medium reactor, respectively. The volume flow ratio of saturated hydrogen, initial hydrogen, reaction hydrogen, and delayed hydrogen is 1:5:12:3. The hydrogenation reaction is carried out in the medium reactor. The total hydrogenation reaction time is 330 min. The residence time of the material in the initial zone, reaction zone, and delay zone accounts for 12.5%, 55%, and 32.5%, respectively. The hydrogenated slurry is then obtained from the top outlet of the medium reactor.

[0080] (3) The hydrogenated slurry and 185°C quench oil are mixed in the feed pipe at a volumetric flow rate ratio of 11:1. The mixture is then introduced into a high-pressure momentum separator at the middle inlet of the separator. The height-to-diameter ratio of the high-pressure momentum separator is 7:2, and the feed level is 50±5%. Simultaneously, 395°C turbulent hydrogen is introduced into the lower part of the high-pressure momentum separator at a volumetric flow rate ratio of 5:3 to the hydrogenated slurry for separation. The separation pressure is 15.6 MPa, and the temperature is 412°C, resulting in a hydrogen-rich oil-gas and solid-liquid phase mixed slurry. After cooling the solid-liquid phase mixed slurry, it is introduced into a low-pressure momentum separator at a level of 55±5% for separation. The separation pressure is 0.45 MPa, and the temperature is 375°C, resulting in heavy components at the bottom of the low-pressure momentum separator. The heavy components are fed into a flash evaporator for flash evaporation at a temperature of 373°C and a pressure of 0.22 MPa. The material exiting from the bottom outlet of the flash evaporator enters a hydrocyclone separator for separation at a temperature of 370°C and a pressure of 1.15 MPa. A supernatant is obtained at the top of the hydrocyclone separator, and a heavy slurry is obtained at the bottom. The heavy slurry is used as a circulating slurry and fed into a deep activation reactor for recycling.

[0081] (4) The supernatant is fed into the middle section of the solid-liquid separation depressurization tower. At the same time, the oil washing tower separates the light components in the low-pressure inertial momentum separator and the flash tower. The temperature of the oil washing tower is 266℃ and the pressure is 0.17MPa. The heavy components obtained after separation are fed into the middle section of the solid-liquid depressurization tower through the bottom outlet of the oil washing tower. Depressurization separation is carried out in the solid-liquid depressurization tower (from top to bottom, the upper narrow tower section, the middle wide tower section and the lower conical section, the upper narrow tower section has diameters of 7m, 10m and 5m respectively). The temperature at the top of the tower is 55℃, the temperature at the bottom of the tower is 312℃, and the tower pressure is -0.07MPa. The residence times of the middle tower section and the lower tower section are 270s and 90s respectively. The total residence time of the middle tower section and the lower tower section is 6.0min. Two side streams are extracted. Diesel products are extracted from the upper section and wax oil products are extracted from the middle section. In addition, heavy wax oil is obtained at the bottom of the solid-liquid vacuum separation tower. The heavy wax oil is used as a circulating slurry and is fed into the deep activation reactor for recycling.

[0082] Example 2

[0083] The difference from Example 1 is that the temperature settings for the various hydrogens in the reactor are different.

[0084] The specific steps for distinguishing (2) are as follows:

[0085] (2) The slurry and 391℃ saturated hydrogen are mixed in the feed pipe. The volume flow ratio of the slurry to the saturated hydrogen is 15:11. The mixture is then introduced into the medium reactor through the bottom inlet of the reactor. The medium reactor is divided into an initial zone for introducing 491℃ initial hydrogen, a reaction zone for introducing 491℃ reaction hydrogen, and a delay zone for introducing 320℃ delayed hydrogen from bottom to top. The initial zone, reaction zone, and delay zone account for 10%, 55%, and 35% of the volume of the medium reactor, respectively. The volume flow ratio of saturated hydrogen, initial hydrogen, reaction hydrogen, and delayed hydrogen is 1:5:12:3. The hydrogenation reaction is carried out in the medium reactor. The total hydrogenation reaction time is 330 min. The residence time of the material in the initial zone, reaction zone, and delay zone accounts for 12.5%, 55%, and 32.5%, respectively. The hydrogenated slurry is then obtained from the top outlet of the medium reactor.

[0086] Comparative Example 1

[0087] The difference from Example 1 is that the temperature, flow rate, and hydrogenation reaction time of various hydrogens in the reactor are set differently, and all are outside the control range.

[0088] The specific steps for distinguishing (2) are as follows:

[0089] (2) The slurry and 385℃ saturated hydrogen are mixed in the feed pipe. The volume flow ratio of the slurry to the saturated hydrogen is 15:11. The mixture is then introduced into the medium reactor through the bottom inlet of the reactor. The medium reactor is configured from bottom to top with an initial zone for introducing 490℃ initial hydrogen, a reaction zone for introducing 490℃ reaction hydrogen, and a delay zone for introducing 280℃ delayed hydrogen. The initial zone, reaction zone, and delay zone account for 10%, 55%, and 35% of the volume of the medium reactor, respectively. The volume flow ratio of saturated hydrogen, initial hydrogen, reaction hydrogen, and delayed hydrogen is 1:7:17:2. The hydrogenation reaction is carried out in the medium reactor. The total hydrogenation reaction time is 380 min. The residence time of the material in the initial zone, reaction zone, and delay zone accounts for 12.5%, 55%, and 32.5%, respectively. The hydrogenated slurry is then obtained from the top outlet of the medium reactor.

[0090] Comparative Example 2

[0091] The difference from Example 1 is that no delayed hydrogen is introduced into the reactor.

[0092] The specific steps for distinguishing (2) are as follows:

[0093] (2) The slurry and saturated hydrogen at 391°C are mixed in the feed pipe. The volume flow ratio of the slurry to the saturated hydrogen is 15:11. The mixture is then introduced into the medium reactor through the bottom inlet of the reactor. The medium reactor is divided into an initial zone for introducing initial hydrogen at 501°C and a reaction zone for introducing reactive hydrogen at 501°C from bottom to top. The initial zone and the reaction zone account for 10% and 90% of the volume of the medium reactor, respectively. The volume flow ratio of saturated hydrogen, initial hydrogen, and reactive hydrogen is 1:5:12. The hydrogenation reaction is carried out in the medium reactor. The total hydrogenation reaction time is 330 min. The residence time of the material in the initial zone and the reaction zone accounts for 12.5% ​​and 87.5%, respectively. The hydrogenated slurry is then obtained from the top outlet of the medium reactor.

[0094] Comparative Example 3

[0095] The difference from Example 1 is that the initial reaction device for the raw oil uses a conventional stirring type for mixing, and hydrogen-rich acidic gas is introduced into the top of the tank.

[0096] The specific steps for distinguishing (1) are as follows:

[0097] (1) After preheating the heavy and inferior oil in a stirred reactor to a temperature of 356°C, hydrogen-rich acid gas (including hydrogen sulfide and hydrogen) is introduced into the top of the tank for hydrogen decomposition. The volume flow ratio of hydrogen-rich acid gas to heavy and inferior oil is 100:675. Then, catalyst precursor (molybdenum isooctanoate) and circulating slurry (including polycyclic aromatic hydrocarbons and catalyst) are introduced into the top of the tank. The mass ratio of catalyst precursor to circulating slurry is 1:700, and the mass ratio of catalyst (molybdenum disulfide) in the slurry is 2300-2500 ppm. After that, the mixture is stirred and dispersed into a slurry to obtain the slurry.

[0098] Comparative Example 4

[0099] The difference from Example 1 is that: no circulating oil washing to stop the reaction process and a cyclone separator are set up, a single slurry balance flash distillation tower is used, and a conventional equal-diameter vacuum tower is used for solid-liquid separation.

[0100] The specific steps for distinguishing between (3) and (4) are as follows:

[0101] (3) The hydrogenated slurry is fed into a high-pressure momentum separator at the middle inlet of the separator. The height-to-diameter ratio of the high-pressure momentum separator is 7:2, and the feed level is 50±5%. Simultaneously, 395℃ turbulent hydrogen is introduced into the lower part of the high-pressure momentum separator. The volumetric flow ratio of the turbulent hydrogen to the hydrogenated slurry is 5:3 for separation. The separation pressure is 15.6 MPa, and the temperature is 412℃, resulting in a hydrogen-rich oil-gas and solid-liquid mixed slurry. After cooling the solid-liquid mixed slurry, it is fed into a low-pressure momentum separator at a level of 55±5% for separation. The separation pressure is 0.45 MPa, and the temperature is 375℃, resulting in heavy components at the bottom of the low-pressure momentum separator. The heavy components are then fed into a flash distillation tower for flash distillation at a temperature of 373℃ and a pressure of 0.22 MPa.

[0102] (4) The material exiting from the bottom outlet of the flash distillation tower is directly fed into the middle section of the solid-liquid separation vacuum distillation tower. Simultaneously, the oil washing tower separates the light components from the low-pressure inertial momentum separator and the flash distillation tower. The temperature of the oil washing tower is 266℃ and the pressure is 0.17MPa. The heavy components obtained after separation are fed into the middle section of the solid-liquid vacuum distillation tower from the bottom outlet of the oil washing tower. Vacuum separation is carried out in the solid-liquid vacuum distillation tower (a conventional equal-diameter vacuum distillation tower with a diameter of 8m). The temperature at the top of the tower is 55℃, the temperature at the bottom of the tower is 312℃, and the pressure is -0.07MPa. The residence times in the middle section and the lower section are 270s and 90s, respectively, and the total residence time in the middle section and the lower section is 6.0min. Two side streams are collected: diesel product is collected from the top section, and wax oil product is collected from the middle section. In addition, heavy wax oil is obtained at the bottom of the solid-liquid vacuum distillation tower. The heavy wax oil is used as a circulating slurry and fed into the deep activation reactor for recycling.

[0103] Table 1

[0104] Case Average yield of wax oil diesel average yield methane volume content Runtime Example 1 35％ 42％ benchmark value 36 months Example 2 32％ 40％ (+)3-6％ 30 months Comparative Example 1 25％ 33％ (+)8-12％ 20 months Comparative Example 2 21％ 31％ (+)10-15％ 18 months Comparative Example 3 26％ 36％ (+)8-12％ 22 months Comparative Example 4 19％ 37％ (+)15-20％ 12 months

[0105] Note: "+" indicates the percentage increase relative to the baseline value; the operating cycle is: the system runs continuously from the start of feeding until the system cannot operate normally or produce qualified products due to excessive pressure drop, which is one cycle; the actual time is the time from system start to stop.

[0106] As shown in Table 1, the wax oil and diesel oil yields of Examples 1-2 were higher than those of Comparative Examples 1-4. Using the methane production of Example 1 as a baseline, the methane production of Comparative Examples 1-4 was significantly higher than that of Example 1, and the increase in methane production in Comparative Examples 1-4 was also significantly higher than that of Example 2. The amount of methane produced reflects the amount of coking in the system. Furthermore, the operating cycle of Examples 1-2 was significantly longer than that of Comparative Examples 1-4.

[0107] Therefore, the system of this invention exhibits high oil yield, high light content, low methane content, and long operating cycle in the hydrogenation of heavy and inferior oil. This invention, through hydrogen decomposition, activation, and mixing and dispersion into a slurry, yields a highly dispersed slurry with more uniform dispersion and less susceptibility to coking. By controlling the temperature and flow rate of hydrogen in the reactor and setting up a high-pressure circulating oil washing process to stop the reaction, targeted conversion of supramolecular asphaltene and residual oil is achieved (products with high oil content, high light content, and low methane). Simultaneously, the generation of suspended coke and wall-phase coke within the reactor is reduced, fundamentally decreasing the formation and polymerization of subsequent coke. By setting up a hot-cut dual-slurry balance flash distillation tower, a cyclone separator, and a multi-stage variable-diameter solid-liquid separation decompression tower, the concentration of small coke lumps is reduced, avoiding an environment conducive to the accumulation and enrichment of coke lumps, and reducing the possibility of already generated coke adhering, enriching, and depositing to form coke lumps. Ultimately, this solves the problem of full-load, long-cycle system operation.

[0108] The above are merely preferred embodiments of the present invention and do not limit the patent scope of the present invention. Any equivalent structural or procedural transformations made using the present invention specification, or direct or indirect applications in other related technical fields, are similarly included within the patent protection scope of the present invention.

Claims

1. A method for hydrogenating heavy, low-quality oil using mesoscale reaction, characterized in that, Includes the following steps: (1) After preheating the heavy and inferior oil, hydrogen-rich acid gas is introduced for hydrogen decomposition. The hydrogen-rich acid gas includes hydrogen sulfide and hydrogen. The volume flow ratio of the hydrogen-rich acid gas to the heavy and inferior oil is 100:650-700. Then, a catalyst precursor is introduced for activation and dispersion. The catalyst precursor is molybdenum isooctanoate to obtain a slurry. The amount of catalyst precursor introduced is such that the mass ratio of the catalyst in the slurry of step (1) is 2100-4500 ppm. (2) After mixing the slurry with saturated hydrogen at 385-395℃ in the feed pipe, it is introduced into the medium reactor from the bottom. The medium reactor is divided into an initial zone for introducing initial hydrogen at 495-505℃, a reaction zone for introducing reactive hydrogen at 495-505℃, and a delay zone for introducing delayed hydrogen at 300-320℃ from bottom to top. The volume flow ratio of saturated hydrogen, initial hydrogen, reactive hydrogen, and delayed hydrogen is 1:4-6:11-13:2-4. The hydrogenation reaction is carried out, and the total time of the hydrogenation reaction is 300-360 min. The residence time in the initial zone, reaction zone, and delay zone is 10-15%, 50-60%, and 30-35%, respectively, to obtain hydrogenated slurry. (3) After slurry separation, the hydrogenated slurry is used to obtain the supernatant; (4) After the supernatant is separated by depressurization, diesel oil and wax oil are obtained respectively.

2. The method for hydrogenating heavy, inferior oil using mesoscale reaction according to claim 1, characterized in that, In step (1), the preheating temperature is 350-360℃.

3. The method for hydrogenating heavy, low-quality oil using mesoscale reaction according to claim 1, characterized in that, In step (2), the volume flow ratio of the slurry to saturated hydrogen is 15:10-12.

4. The method for hydrogenating heavy, inferior oil using mesoscale reaction according to claim 1, characterized in that, In step (2), the initial zone, reaction zone, and delay zone occupy 10-15%, 50-60%, and 30-35% of the volume of the reactor, respectively.

5. The method for hydrogenating heavy, inferior oil using mesoscale reaction according to any one of claims 1-4, characterized in that, In step (3), the slurry separation includes high-pressure momentum separation and low-pressure momentum separation performed sequentially.

6. The method for hydrogenating heavy, inferior oil using mesoscale reaction according to claim 5, characterized in that, The high-pressure momentum separation includes: passing hydrogenated slurry and quenched oil at 175-190℃ into a high-pressure momentum separator, then introducing disturbing hydrogen into the lower part of the separator for separation to obtain hydrogen-rich oil gas and solid-liquid mixed slurry.

7. The method for hydrogenating heavy, inferior oil using mesoscale reaction according to claim 6, characterized in that, The low-pressure momentum separation includes: passing a solid-liquid phase mixed slurry into a low-pressure momentum separator, and obtaining heavy components after separation; and obtaining supernatant by flash evaporation and cyclone separation of the heavy components.

8. The method for hydrogenating heavy, inferior oil using mesoscale reaction according to claim 1, characterized in that, In step (4), the pressure reduction separation is carried out in a solid-liquid pressure reduction separation tower, which consists of an upper narrow tower section, a middle wide tower section and a lower conical section from top to bottom; the tower top temperature is 52-62℃, the tower bottom temperature is 305-315℃, and the tower pressure is -0.075MPa to -0.065MPa. Two side streams are used for extraction, with diesel products extracted from the upper section and wax oil products extracted from the middle section.

9. The method for hydrogenating heavy, inferior oil using mesoscale reaction according to claim 7, characterized in that, The heavy slurry obtained after cyclone separation and the heavy slurry obtained after depressurization separation in step (4) are returned to step (1) as circulating slurry and are introduced together with the catalyst precursor for activation and dispersion.

10. The method for hydrogenating heavy, low-quality oil using mesoscale reaction according to claim 1, characterized in that, The system for implementing the method includes a feedstock initial reaction device, a medium reactor, a slurry separation device, and a vacuum separation device connected in sequence. The feedstock initial reaction device includes a desiccant preheater, a hydrogen desiccant reactor, and a deep activation reactor connected in sequence. The medium reactor has a slurry inlet at the bottom and a slurry outlet at the top. The slurry inlet is connected to the deep activation reactor through a feed pipe, which also has a saturated hydrogen inlet. The medium reactor includes an initial zone, a reaction zone, and a delay zone arranged from bottom to top. The initial zone has an initial hydrogen inlet, the reaction zone has a reaction hydrogen inlet, and the delay zone has a delay hydrogen inlet.

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