WO2011047538A1 - 一种沸腾床反应器及其加氢方法 - Google Patents
一种沸腾床反应器及其加氢方法 Download PDFInfo
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- WO2011047538A1 WO2011047538A1 PCT/CN2010/001641 CN2010001641W WO2011047538A1 WO 2011047538 A1 WO2011047538 A1 WO 2011047538A1 CN 2010001641 W CN2010001641 W CN 2010001641W WO 2011047538 A1 WO2011047538 A1 WO 2011047538A1
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- Prior art keywords
- bed reactor
- catalyst
- ebullated
- reactor
- reactor according
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- C—CHEMISTRY; METALLURGY
- C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
- C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
- C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
- C10G2300/10—Feedstock materials
- C10G2300/107—Atmospheric residues having a boiling point of at least about 538 °C
-
- C—CHEMISTRY; METALLURGY
- C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
- C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
- C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
- C10G2300/10—Feedstock materials
- C10G2300/1077—Vacuum residues
-
- C—CHEMISTRY; METALLURGY
- C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
- C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
- C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
- C10G2300/40—Characteristics of the process deviating from typical ways of processing
- C10G2300/4018—Spatial velocity, e.g. LHSV, WHSV
Definitions
- the present invention relates to a reactor and a process for its use, and in particular to an ebullated bed reactor and a hydrogenation process therefor. Background technique
- Heavy oil processing technology mainly includes two processes of decarburization and hydrogenation:
- the decarburization process mainly includes solvent deasphalting, coking, heavy oil catalytic cracking, etc.
- the decarbonization process has low investment in equipment, the liquid product yield is low and the quality is poor, which cannot meet the current environmental protection requirements.
- the proportion of residual oil to crude oil increases year by year, up to 70wt%, and the most commonly used decarbonization process of heavy and residual oil is coking.
- the by-product of the process is a large amount of low value added coke.
- the hydrogenation process can be specifically classified into fixed bed hydrogenation, moving bed hydrogenation, suspension bed hydrogenation, and bubbling bed hydrogenation depending on the state of presence of the catalyst in the reactor.
- the investment is high due to the use of high-pressure reaction equipment, but the product quality is good and the liquid yield is high, so that the weight and the residue can be minimized and lightened.
- the relatively mature residue hydrogenation process is a fixed bed residue oil hydrogenation, but the process is restricted by the nature of the raw materials, and the requirements for the metal and residual carbon of the raw materials are relatively strict.
- the suspended bed and moving bed technologies have certain advantages in heavy oil processing, but they have developed slowly in recent years.
- the moving bed hydrogenation technology usually adopts the operation mode of feedstock oil and catalyst countercurrent or cocurrent flow through the reactor, and uses the initial activity of the catalyst to carry out heavy oil processing. Although the hydrogenation effect is better, the catalyst is used in a large amount, and the hydrogenation activity of the catalyst is not Being fully utilized.
- the bubbling bed reactor is a gas, liquid and solid three-phase fluidized bed. It can process heavy and inferior feedstock oil with high metal and high asphaltene content. It has small pressure drop and uniform temperature distribution, and can maintain catalyst activity throughout the operating cycle. Constant, can be run It is characterized by the addition of fresh catalyst and the removal of spent catalyst.
- the on-line addition and discharge technology of the catalyst is the key technology to ensure the quality of the fluidized bed, stable operating conditions and operating cycle.
- the catalyst in-line addition mode of the fluidized bed hydrogenation technology is usually carried out by gas phase transportation, liquid phase transportation method or by gravity to directly add the solid catalyst from the high pressure storage tank located at the upper portion of the reactor to the bubbling bed reactor.
- the catalyst In the bubbling bed reactor, the catalyst is kept in a good boiling state, and the liquid viscosity, reaction pressure, gas-liquid flow rate and reaction temperature in the reactor are kept constant, and the fresh catalyst is directly added to the bubbling bed reactor.
- the instantaneous fluctuation of the above conditions causes the instantaneous operation of the fluid state and operating conditions in the reactor to be unstable.
- the fresh catalyst since the fresh catalyst has high initial activity, it is directly added to the bubbling bed reactor, and the inferior heavy and residual oil raw materials are used. Contact mixing will result in rapid carbon deposition of the catalyst, rapid loss of activity, affecting the hydrogenation reaction of the reactant stream, and increasing the replacement frequency of the catalyst.
- CN101418222A, CN1335357A, CN101360808A are prior art methods for treating inferior residue oil, wherein CN101418222A adopts a combined reaction device of a fluidized bed and a suspended bed; CN1335357A adopts a combined reaction device of an expanded bed and a moving bed; CN101360808A includes at least Two series of upflow reactors, but none of these prior art techniques disclose an on-line process for the catalyst when the catalyst in the reactor fails to achieve the desired activity.
- U.S. Patent 4,398,852 describes an on-line addition of a catalyst for an ebullated bed reactor.
- the method comprises first adding a catalyst to a vessel of a high pressure resistant catalyst, then charging it to a reaction pressure, and then opening a valve located on the catalyst vessel and the reactor connecting line to cause the catalyst to enter the bubbling bed reactor by gravity. .
- the process is to directly add the catalyst to the bubbling bed reactor by gravity.
- the catalyst is added in such a way that the initial active catalyst rapidly deposits carbon when it contacts the inferior raw material, accelerates the deactivation rate, and increases the replacement frequency of the catalyst. Because the preheating temperature of the catalyst and hydrogen is lower than the reaction temperature, the boiling bed reaction temperature fluctuates, the operating conditions are not stable, and the product quality is affected.
- a typical bubbling bed process is described in US Re 25 770 and U.S. Patent 4,398, 852.
- An internal circulation cup is provided in an ebullated bed reactor for gas-liquid separation to increase liquid conversion.
- the process has the following disadvantages in practical applications: less catalyst inventory in the reactor, low reactor space utilization; high maintenance cost of the circulating oil pump, and once the circulating oil pump malfunctions or is damaged, it will cause the catalyst to sink and collect.
- the device is forced to stop running; the liquid product in the reactor stays for too long under non-catalytic hydrogenation conditions, and it is easy to carry out secondary thermal cracking coking at high temperature to reduce product quality.
- the patents CN02109404.7 and CN101376092A respectively introduce a novel bubbling bed reactor, which uses a three-phase separator and a three-phase separator with a pilot port for efficient gas, liquid and solid separation. Compared with a typical ebullated bed reactor, it has a simple structure, easy operation and high reactor utilization. Features. However, because the high-diameter of the bubbling bed reactor is relatively large, usually 1:6 to 1:8, and most of the effective reaction space in the reactor except the three-phase separator at the top is an empty cylinder structure, lacking forced mass transfer. The structure, so the mass transfer effect between gas, liquid and solid is poor, the hydrogenation effect of the liquid phase product is not significant, and the product quality is poor. In addition, the bubbling bed reactor is a back-mixing reactor, and some of the raw materials that are expected to be reacted in the future will be discharged from the reactor along with the post-reaction stream, so that the conversion rate of the raw materials is relatively low.
- the prior art there is a bubbling bed containing two or more reaction stages, and a three-phase separation component is arranged in the bubbling bed reactor for gas, liquid and solid separation, and at the same time, hydrodemetallization, hydrodesulfurization, hydrodenitrogenation are realized.
- the sequence of the plurality of reactions is arranged, wherein one to two catalysts are used per reaction section.
- the three-phase separating component is composed of a flow guiding member and a baffle member, wherein the guiding member is a cone or a cone having an upper narrow width or an upper width and a lower width, and has a lower position and a lower position, and a lower position guide.
- the reactor structure is actually a large reactor in which two reactors are combined, the pipeline between the reactors and other equipment such as separators, sedimentation tanks and the like are removed, and the advantage is that the heat energy can be reasonably utilized, but there is still
- the reactor is bulky, which increases the difficulty of transportation, installation, daily operation and maintenance; in a reactor, as the reaction section increases, the number of three-phase separators increases, making the entire reactor structure complicated, and The more the three-phase separator occupies more reactor space, the smaller the effective reaction space for gas, liquid and solid, and the higher the cost of high temperature and high pressure ebullated bed reactor, which will result in unreasonable utilization of resources; Operation mode, high requirements for operation and gas-liquid distribution plate between sections, requiring very stable operating conditions. If the device fluctuates instantaneously and will affect the separation effect of the three-phase separator, the solids entrained by the gas and liquid can be
- the present invention provides an ebullated bed reactor.
- a number of internal circulation zones are added to the fluidized bed reactor to effectively increase the conversion rate of raw materials.
- An ebullated bed reactor comprising a reactor housing perpendicular to the ground, a phase separator located in an upper portion of the housing, and an inner circulation zone disposed below the phase separator, the inner circulation zone including a cylinder and a cone a shaped diffusion section and a flow guiding structure, the cylinder and a conical diffusion section disposed at a lower end of the cylinder are located in the reactor housing, and the diversion structure is located on the inner wall of the reactor housing at the lower end of the conical diffusion section
- the flow guiding structure is an annular convex structure disposed on the inner wall of the reactor, and the longitudinal section along the axis of the reactor is one of a trapezoidal shape, an arcuate shape, a semicircular shape, a triangular shape, or other equivalent replacement which can serve as a guide.
- the improved shape should be included in the scope of the present invention.
- the upper end of the cylinder is a slightly outwardly diffusing horn Mouth structure.
- the specific structure of the novel bubbling bed reactor of the present invention is:
- a raw material inlet and a gas-liquid distribution plate are provided at the bottom of the reactor housing.
- a gas outlet is arranged at the top of the reactor shell, and the upper shell wall is provided with a liquid discharge port, and the liquid discharge port is located on the reactor shell wall between the upper end of the inner cylinder of the phase separator and the lower end of the outer cylinder, and is used for The gas and liquid generated by the reaction are led out.
- the phase separator is disposed in an upper space of the casing, and includes two concentric cylinders having different inner diameters: an inner cylinder and an outer cylinder.
- the upper and lower ends of the inner cylinder and the outer cylinder are all open, the upper end opening of the outer cylinder is higher than the upper end opening of the inner cylinder, and the lower end opening of the outer cylinder is also higher than the lower end opening of the inner cylinder.
- the lower end of the inner cylinder is a conical diffusion section, and the opening of the diffusion section (ie, the lower end opening of the inner cylinder) has a diameter smaller than the inner diameter of the reactor; the lower end of the outer cylinder is also a conical diffusion section, and the diffusion section
- the opening i.e., the lower end opening of the outer cylinder
- the inner cylinder of the phase separator constitutes a central tube of the separator
- the annular space between the inner cylinder and the outer cylinder constitutes a baffle of the phase separator
- the annular space between the outer cylinder and the inner wall of the reactor is the phase a clear liquid product collection area of the separator
- the lower end opening of the inner cylinder is a flow introduction port
- the annular opening formed by the lower end opening of the inner cylinder and the inner wall of the reactor is a catalyst discharge port of the phase separator, and the solid separated From this point the particulate catalyst is returned to the catalyst bed.
- each component of the phase separator can be determined by the designer in the art according to the specific requirements of the catalyst size, the reactor treatment amount, the reaction conditions and the separation effect, by calculation or simple experiment. Or, conventional techniques disclosed in the art may be employed, for example, reference may be made to the patent CN02109404.7 or CN101376092A previously filed by the applicant.
- the inner circulation zone includes a cylinder, a tapered diffusion section, and an adjacent flow guiding structure.
- the cylinder is connected to the conical diffusion section.
- the diameter of the lower end opening of the conical diffusion section is smaller than the inner diameter of the reactor, and the conical flow structure adjacent to the conical diffusion section of the cylinder is combined to form an inner circulation zone.
- One or more inner circulation zones may be provided in the reactor according to the requirements of the height to diameter ratio and the conversion depth of the reactor, preferably 2 to 3 inner circulation zones, wherein the inner diameters of the cylinders of different inner circulation zones may be the same or different .
- the flow guiding structure is an annular convex structure disposed on the inner wall of the reactor casing, and the longitudinal section along the axis of the reactor is trapezoidal, arcuate, semi-circular, triangular, or other shape, which can serve as an equivalent replacement for guiding, Improvements and the like should be included in the scope of the present invention.
- the angle formed between the tangential line at the intersection of the flow guiding structure on the side of the phase separator and the reactor wall and the inner wall of the reactor is called the coverage angle, and the coverage angle is an acute angle, preferably less than 60 degrees; on the opposite side, That is, the angle formed by the tangential line at the intersection of the flow guiding structure away from the phase separator and the reactor wall and the inner wall of the reactor is called the friction angle, and the friction angle is also an acute angle, preferably less than 60 degrees.
- the flow guiding structure is surrounded by a straight
- the diameter is between the outer cylinder of the phase separator and the inner cylinder diameter.
- a flow guiding structure may be disposed immediately below the phase separator, the flow guiding structure being located in the upper middle portion of the reactor between the phase separator and the inner circulation zone.
- the flow guiding structure is similar to the flow guiding structure in the inner circulation zone.
- the gas discharge port is generally located at the center of the top of the reactor.
- the liquid discharge port is generally disposed at the upper portion of the reactor wall, and is positioned between the upper end opening and the lower end opening of the phase separator outer cylinder.
- the upper portion of the phase separator is typically provided with a buffer space from which the phase separated gas product is concentrated and exits the reactor from the gas discharge port.
- the reactor has an aspect ratio in the range of 0.01 to 0.1.
- the bubbling bed reactor of the present invention typically also includes at least one component that discharges the catalyst from the reactor, and at least one component that replenishes the reactor with fresh catalyst.
- the components that replenish the fresh catalyst are typically disposed at a top position of the reactor, and the components that discharge the catalyst are typically located near the bottom of the reactor.
- a catalyst addition tube is provided at the top of the reactor housing, and a catalyst discharge tube is provided at the bottom.
- the catalyst replacement system and method of use may be any suitable apparatus or method, such as the method described in U.S. Patent No. 3,398,085 or U.S. Patent 4,398,852.
- a gas-liquid distribution plate should generally be disposed at the bottom of the cylindrical reactor casing, and the gas-liquid distribution plate can be selected to uniformly distribute the gas and the liquid phase.
- a bubble cap structure can be employed.
- the internal circulation zone of a fluidized bed reactor works as follows: The flow passes through different reactor cross-sectional areas and the flow rate changes.
- the stream in the bubbling bed reactor consists of gas, liquid, and solid three phases, namely a solid catalyst, a liquid reactant stream, gaseous hydrogen, and light hydrocarbons formed.
- the cross-sectional area of the stream passing through the reactor changes, the gas and liquid flow rates change accordingly, and the catalyst entrained by the gas-liquid will rapidly rise or settle; the liquid phase feedstock reacts in the ebullated-bed reactor.
- the process proceeds to form a part of the light component, which will pass the hydrogen up through the reactor, and the partially reacted liquid phase product and the unreacted raw material will be similar to the movement state of the catalyst, and the cross section in the reactor is small.
- the fluid acceleration zone flows rapidly upwards, and a countercurrent phenomenon opposite to the flow direction of the main stream occurs at an instantaneous expansion of the cross-sectional area.
- the use of an ebullated-bed reactor with an internal circulation zone and a phase separator for hydrotreating inferior feedstock oil can increase the conversion of liquid phase heavy components, and the structure of the fluidized bed reactor can improve the flow between the reactants in the reactor. Quality and heat transfer effects.
- Another object of the present invention is to provide a hydrogenation process for an ebullated bed reactor, which is capable of The boiling bed device is operated smoothly when the catalyst is added, thereby ensuring the operation cycle of the device, and further processing the fluid after the fluidized bed reaction, thereby improving product quality.
- the expanded bed reactor is connected to the bubbling bed reactor through a pipeline; when the catalyst in the bubbling bed reactor does not meet the catalytic activity requirement, when the product quality cannot be ensured, a fresh catalyst needs to be replenished, and the required catalyst is expanded.
- the bed reactor is replenished and the catalyst lacking in the expanded bed reactor is supplemented by a fresh catalyst addition tank.
- the expanded bed reactor described in the process of the present invention has a bed expansion ratio of 5 v% to 25 v%, preferably 10 v% to 25 v%, and most preferably 15 v% to 20 v%.
- the expansion ratio described in the present invention means the ratio of the height of the surface after expansion of the catalyst to the height of the surface after the natural filling of the catalyst is the same as the height of the surface after the natural filling of the catalyst.
- the operating conditions of the expanded bed reactor are: a reaction pressure of 6 to 30 MPa, preferably 10 to 18 MPa; a reaction temperature of 350 to 500 ° C, preferably 380 to 430 ° C; and a space velocity of Ol ⁇ h' 1 , Preferably, it is ⁇ ⁇ 1 ; the hydrogen oil volume ratio is from 400 to 2,000, preferably from 600 to 1,500.
- the amount of catalyst added per time of the expanded bed reactor is 2 to 20 times that of the in-line addition of the catalyst in the bubbling bed reactor.
- the amount of catalyst in the expanded bed reactor is the amount of on-line displacement of the single-boiler reactor When it is ⁇ 5 times, it is replenished from a catalyst addition tank provided in the upper portion of the reactor.
- the proportion of the liquid phase product after the gas-liquid separation into the expanded bed reactor is from 5 wt% to 70 wt%, preferably from 10 wt% to 50 wt%, based on the total amount of the liquid phase product after the reaction.
- the inferior raw material oil described in the method of the present invention comprises one or more of inferior raw materials such as atmospheric residue, vacuum residue, deasphalted oil, oil sand bitumen, thick crude oil, coal tar and coal liquefied heavy oil.
- the operating conditions of the bubbling bed reactor are: a reaction pressure of 6 to 30 MPa, preferably 10 to 18 MPa; a reaction temperature of 350 to 500 ° C, preferably 400 to 450 ° C; a space velocity of 0.1 to 5 °, preferably It is 0.5 ⁇ 3h; the hydrogen oil volume ratio is 400 ⁇ 2000, preferably 600 ⁇ 1500.
- the catalyst used in the above reactor may be a conventional bubbling bed hydrogenation catalyst in the art, and the properties of a typical catalyst are: refractory inorganic oxide as a carrier, and VIB
- the family and/or the third group of metals are active components.
- the catalyst has a particle diameter of 0.8 mm and a particle length of 3 to 5 mm.
- the basic physicochemical properties are substantially the same as those of the conventional fixed bed hydrogenation catalyst.
- the present invention preferably uses a catalyst having the following properties:
- the catalyst particles have a diameter of from 0.1 to 0.8 mm, preferably from 0.1 to 0.4 mm, and the catalyst contains a Group IB and a Group of the active hydrogenation metal component.
- the carrier is A1 2 0 3 and contains at least one auxiliary agent, and the auxiliary agent is selected from the following elements: B, Ca, F, Mg, P, Si, Ti, etc.
- the additive content is from 0,5 wt% to 5.0 wt%.
- the pore volume of the catalyst is 0.6 ⁇ 1.2mL/g, the pore volume with pore diameter ⁇ 8nm is less than 0.03mL/g, generally 0.005 ⁇ 0.02mL/g, the average pore diameter is 15 ⁇ 30nm, and the pore diameter is between 15 ⁇ 30nm. It accounts for more than 50% of the total pore volume, generally 50% ⁇ 70%.
- the specific surface is 100 to 300 m 2 /g, preferably 120 to 240 m 2 /g.
- the catalyst comprises a Group VIB metal oxide (such as Mo0 3 ) 1.0 wt% to 20.0 wt%, preferably 3.0 wt% to 15.0 wt%, and contains a Dishan metal oxide (such as MO or CoO) 0.1 wt% to 8.0 wt%. It is preferably from 0.5% by weight to 5.0% by weight. Catalyst wear ⁇ 0.1 wt%.
- a Group VIB metal oxide such as Mo0 3
- Dishan metal oxide such as MO or CoO
- the catalyst used in the bubbling bed reactor is a microsphere catalyst having a particle size of between 0.1 and 0.8 mm, and the conventional circulating bed reactor, such as the internal circulation set in the reactor disclosed in US Re 25570 and related patents.
- the main function of the cup is to perform efficient separation of gas and liquid.
- the catalyst used in the reactor is substantially the same size as the conventional hydrogenation catalyst, so the conventional bubbling bed reactor is not suitable for the method of the present invention.
- the bubbling bed reactor used in the process of the present invention may be selected from an ebullated-bed reactor in which internal components such as a gas, a liquid, a solid three-phase separator, and a gas inlet are provided. Both the bubbling bed reactors disclosed in CN1448212A and CN101376092A can be used in the hydrogenation process of the present invention.
- the above reactor can be used in the method of the present invention, since most of the effective reaction zone of the bubbling bed reactor is an empty cylinder structure, the mass transfer effect between the gas-liquid solid is poor, so the hydrogenation effect is not significant, and The reactor in the backmixing operation will cause some of the feed that will be reacted in the future to flow out of the reactor as the reaction proceeds, and the conversion rate of the raw materials is relatively low.
- the inferior feedstock oil hydrotreating method of the invention adopts a combined process of a fluidized bed and an expanded bed with an inner circulation zone and a three-phase separator reactor to treat heavy oil raw materials, thereby improving the quality of the light oil product and ensuring the main reactor
- the bubbling bed reactor operates smoothly during the catalyst addition process.
- the combined process can perform heavy oil processing in a flexible operation mode.
- the heavy oil feedstock is subjected to a hydrocracking reaction through a fluidized bed hydrogenation reactor, and the reactants are passed to a separation device to obtain a gas phase product and a liquid phase product, wherein a part of the liquid phase product is recycled. Further hydrotreating to the expanded bed reactor.
- the expanded bed reactor stream entrains the catalyst from the lower portion of the reactor to the bubbling bed reactor.
- This flexible operation overcomes the reactor temperature and pressure fluctuations that may be caused by the direct addition of existing catalysts to the ebullated-bed reactor, eliminating the effects of existing dosing methods on catalyst boiling and reaction stream properties (eg, causing operations) Unstable, such as catalyst entrainment or catalyst bed expansion is not ideal, affecting product quality and equipment operating cycle, etc.).
- the fresh catalyst is first added to the expanded bed reactor, and then added to the bubbling bed reactor, which can serve as a buffering effect, and the stream can be preheated to the reaction temperature in the expanded bed reactor to enter the subsequent bubbling bed reaction.
- the temperature of the liquid, catalyst and gas is substantially equal to the reaction temperature of the bubbling bed to ensure smooth operation of the fluidized bed unit.
- the fresh catalyst is first contacted with the highly improved stream after hydrogenation of the bubbling bed, and the catalyst can be fully utilized.
- the initial activity of the chemical agent avoids the excessive carbon deposition in the initial stage of the catalyst and affects the effect of the catalyst.
- the formation of one or more recycle zones in the bubbling bed reactor provided in the process of the present invention can form a plurality of boiling operating zones, making the overall bubbling bed reactor more flexible to operate.
- the bubbling bed reactor is a reactor with a high degree of backmixing
- the reactor effluent contains a part of unconverted raw materials, and a plurality of small inner circulation zones are provided, which can cause multiple conversion of the raw materials, which is beneficial to Increase conversion rates.
- the phase separator of the fluidized bed reactor has a cylindrical simple structure, and has the advantages of simple preparation process, low production cost, and easy installation and maintenance as compared with the prior art cone or cone phase separator.
- the fresh catalyst enters the expanded bed reactor and then enters the bubbling bed reactor, which acts as a buffer and preheats the catalyst, which makes the operation of the whole operation cycle stable, and at the same time makes the fresh catalyst and the fluidized bed with greatly improved properties.
- the contact of the liquid product can make full use of the initial activity of the catalyst and ensure that the activity of the catalyst can be stably exhibited.
- FIG. 1 is a schematic view of the structure of a fluidized bed reactor of the present invention (only one inner circulation zone is provided).
- FIG. 2 is a schematic flow chart of a method for hydrogenating a fluidized bed reactor of the present invention.
- the reference numerals in the embodiments of the present invention are as follows: 1 catalyst, 2 - catalyst storage tank, 3 - expanded bed reactor, 4, 5, 11 - hydrogen, 6 - heavy hydrocarbon feedstock, 7 - bubbling bed reactor, 8 - high pressure separation device, 9 - cooling and purifying device, 10—distillation unit, 12-gasoline, 13-diesel, 14-hydrogenated tail oil, 15—catalyst discharge line, 16, 17, 18 one valve;
- 101 feed inlet
- 102 gas-liquid distributor
- 103 reactor housing
- 104 diversion structure
- 105 one inner cylinder
- 106 outer cylinder
- 107 catalyst bed
- 108 inlet
- 109 Catalyst inlet
- 111 one phase separator
- 112 - liquid discharge port 112 - liquid discharge port
- 113 - discharge port 112 - discharge port
- 114 - cylinder 115 - conical diffusion section
- 116 catalyst discharge port.
- the bubbling bed reactor used comprises an inner circulation zone and a three-phase separator, the structural features and working principle of which are:
- the catalyst capacity in the reactor shell 103 is at least 35% of the reactor volume, usually 40. % ⁇ 70%, preferably 50% ⁇ 60%.
- the catalyst bed expands to a certain height, and its expanded volume is usually 20% to 70% larger than its static volume.
- the gas-liquid stream entering the reaction zone is then contacted with the catalyst, and the reacted gas-liquid stream and the unconverted feedstock and hydrogen carry the solid catalyst along the axis of the reactor to rise into the flow-conducting structure 104, the cylinder 114 and the cone.
- the circulation zone formed by the shaped diffusion section 115, the flow guiding port 108 formed by the flow guiding structure 104 and the tapered diffusion section 115 of the cylinder are collected into the fluid passage of the cylinder 114, since the cross-sectional area of the fluid passage is reduced at this time. Therefore, the gas-liquid flow rate is increased.
- the fluid passage is instantaneously expanded, so the flow velocity of the gas-liquid fluid is instantaneously reduced, and the ability to carry the solid catalyst is lowered, so that the partially reacted liquid and the unconverted raw material and the solid catalyst are along
- the passage formed by the outer wall of the cylinder and the inner wall of the reactor flows downward into the flow guiding port and mixes with the stream flowing upward from the lower portion of the reactor, thereby forming a small circulation zone.
- the gas phase stream rising from the circulation zone and a portion of the liquid phase stream and the carried catalyst enter the gas separation port 108 enclosed by the flow guiding structure 104 of the phase separator and enter the phase separator 111 for phase separation: the gas is first separated and passed through the gas.
- the discharge port 110 exits the reactor, and the separated catalyst is returned to the reaction zone through the discharge port 113, while the clear liquid phase product substantially free of catalyst particles is discharged from the reactor through the liquid discharge port 112.
- the catalyst can be replenished to the reaction system through the catalyst addition pipe 109 at the upper portion of the reactor, and the portion is passed through the discharge pipe 116 at the lower portion of the reactor.
- the deactivated catalyst exits the reaction system.
- the longitudinal section of the flow guiding structure 104 along the axis of the reactor is trapezoidal, and the coverage angle and the friction angle are both acute angles, preferably less than 60 degrees.
- the longitudinal section of the flow guiding structure 104 along the axis of the reactor can also be arcuate or other suitable shape.
- the phase separator 111 is composed of a concentric cylindrical inner cylinder 105 having an inner diameter, an outer cylinder 106, and an inner wall of the reactor casing 103.
- the inner cylinder 105 constitutes a central pipe of the phase separator, and an annular space between the inner cylinder 105 and the outer cylinder 106 constitutes a baffle of the phase separator, and an annular ring between the outer cylinder 106 and the inner wall of the reactor casing 103
- the space is a clarified liquid product collection zone, and the opening of the lower end of the central pipe is a flow introduction port, and the annular opening formed by the opening of the diffusion section and the inner wall of the reactor casing 103 is a catalyst discharge port.
- the cone apex angle of the outer cylinder diffusion section is generally at least 20 degrees smaller than the cone apex angle of the inner cylinder diffusion section, preferably 40 to 80 degrees.
- the bubbling bed reactor hydrogenation process of the present invention comprises the following steps: The mixture of the heavy hydrocarbon-containing feedstock 6 and the hydrogen gas 5 is heated in a heating furnace, and the above stream is introduced into the bubbling bed reactor 7 to be contacted with the catalyst. After the bubbling bed hydrogenation reaction, the stream is discharged from the top of the reactor into the high pressure separation device 8 for gas-liquid separation, and the partially separated liquid phase product is mixed with the hydrogen gas 4 into the expanded bed reactor 3 for further hydrotreating reaction. The product is discharged from the upper portion of the expanded bed reactor 3 into the high pressure separation unit 8.
- the gas phase can be used as the circulating hydrogen 11, and the condensed light component is mixed with a part of the liquid phase stream of the separation device into the distillation device 10 to obtain the gasoline 12, the diesel 13 And hydrogenated tail oil 14.
- the hydrogenation tail oil can be used as a feedstock for catalytic cracking or fixed bed residue hydrogenation, or recycled back to the ebullated bed reactor.
- catalyst activity in the ebullated-bed reactor is lowered to provide a product of the required quality, catalyst replacement is required. The operation at this time is: discharging the deactivated catalyst in the partial ebullated-bed reactor through the catalyst discharge line 15 to the reactor.
- the valve 18 on the connecting line of the expanded bed reactor 3 and the bubbling bed reactor 7 is then opened, while the valve 17 located in the expanded bed reactor material discharge line is closed, so that the solid catalyst-containing stream enters the bubbling bed reactor 7, online
- the time for adding the catalyst is 10 to 50 minutes.
- Ebullated-bed reactor 7 After the catalyst is added, the normal process is resumed.
- the method of adding the catalyst to the expanded bed reactor 3 is: First, the catalyst 1 is charged into the catalyst storage tank 2, and the storage tank is charged to a pressure slightly higher than the expansion bed pressure by 1 to 5 Pa, and the catalyst storage tank 2 and the expanded bed 3 are opened. Between the valves 16, fresh catalyst is added to the expanded bed reactor 3.
- Example 1 The properties of the residue materials used in the tests are listed in Table 1. It can be seen from Table 1 that the residue has a carbon residue value of 13.6% by weight, a metal content of 141.9 ⁇ m, an asphaltene content of 6.4% by weight, a sulfur content of 2.5% by weight, and a nitrogen content of 0.6% by weight.
- the microspherical catalyst used in the test process has the following properties: an average diameter of 0.6 mm, with alumina as a carrier, a catalyst containing Mo0 3 of 11.2 wt%, a NiO content of 3.0 wt%, a P content of 1.4 wt%, and a pore volume of 0.60 mL. /g, specific surface area is 140m 2 /g, ⁇ 8nm pore volume accounts for 2.6% of total pore volume, and 15 ⁇ 30nm pore volume accounts for 65% of total pore volume.
- the diameter of the cylindrical part of the central tube is 92mm
- the diameter of the bottom of the lower opening of the inner cylinder is 144mm
- the height of the lower part of the inner cylinder is 41mm
- the diameter of the cylindrical part of the outer cylinder is 128mm
- the diameter of the bottom of the tapered part is 138mm.
- the height of the tapered portion is 64 mm
- the upper opening of the outer cylinder is higher than the upper opening of the inner cylinder
- the bottom position of the tapered opening of the lower portion of the outer cylinder is higher than the bottom position of the tapered opening of the lower portion of the inner cylinder, and the height difference between the two is 38 mm.
- the vertical distance between the upper opening of the outer cylinder and the tangent to the top of the reactor shell is 200 mm
- the vertical distance of the center of the liquid product tube from the top tangent of the reactor is 338 mm.
- the annular guide structure has a coverage angle of 20°, a friction angle of 28°, and a guide opening diameter of 100 mm next to the phase separator.
- the diameter of the guide opening forming the circulation zone is 100 mm
- the inner diameter of the cylinder is 80 mm
- the height of the cylinder is 100 mm
- the diameter of the bottom of the tapered diffusion section is 150 mm
- the height of the tapered diffusion section is 45 mm. Comparative example 1
- This example is an example of a hydrogenation process for a fluidized bed reactor of the present invention.
- the schematic of the operation flow is shown in Fig. 2.
- An inner circulation zone is provided in the fluidized bed reactor.
- the process is as follows: After the heavy hydrocarbon-containing feedstock 6 and the hydrogen gas 5 are mixed, the above flow enters the bubbling bed reactor 7 to contact the catalyst. After the bubbling bed hydrogenation reaction, the stream is discharged from the top of the reactor into the high-pressure separation device 8 for gas-liquid separation, and 15% by weight of the liquid phase product of the liquid phase after the reaction is mixed with the hydrogen gas 4 into the expanded bed reactor 3 for further flow. In the hydrotreating reaction, the reaction product is discharged from the upper portion of the reactor into the high pressure separation device 8.
- the gas stream separated by the high pressure separation device is passed through a cooling and purifying device 9 After treatment, the gas phase can be used as recycled hydrogen 11 and the condensed light components are mixed with a portion of the liquid phase stream of the separation unit into distillation unit 10 to provide gasoline 12, diesel 13 and hydrogenated tail oil 14.
- a cooling and purifying device 9 After treatment, the gas phase can be used as recycled hydrogen 11 and the condensed light components are mixed with a portion of the liquid phase stream of the separation unit into distillation unit 10 to provide gasoline 12, diesel 13 and hydrogenated tail oil 14.
- the operation is: discharging the deactivated catalyst in the partial ebullated-bed reactor out of the reactor through the catalyst discharge line 15, and then opening a valve 18 located in the connecting line between the expanded bed reactor 3 and the bubbling bed reactor 7, while closing the valve 17 on the expanded bed reactor material discharge line, so that the solid catalyst-containing stream enters the bubbling bed reactor 7, and the catalyst is added online.
- the time is 20 minutes.
- the catalyst is added in the expanded bed reactor by first loading the catalyst 1 into the catalyst storage tank 2, charging the storage tank to a pressure slightly higher than the expanded bed by 2 Pa, opening between the catalyst storage tank 2 and the expanded bed 3. Valve 16, fresh catalyst is added to the expanded bed reactor.
- the catalyst bed in the expanded bed reactor has an expansion height of 20 v%, and the amount of catalyst added per time in the expanded bed reactor is 10 times that of the in-line addition of each ebullated bed catalyst, when the catalyst remaining in the expanded bed reactor When the amount is 4 times the amount of on-line displacement of a single ebullated-bed reactor, fresh catalyst is added from the catalyst storage tank 2.
- This embodiment is a comparative embodiment of the present invention, and the process is substantially the same as that of Example 2, except that no expanded bed is provided, and when the activity of the catalyst in the ebullated-bed reactor is reduced to a product that cannot provide the index, it is required.
- fresh catalyst is added, fresh catalyst is added directly from the catalyst storage tank disposed at the top of the bubbling bed reactor to the bubbling bed reactor, and the addition process is followed by adding fresh catalyst from the catalyst storage tank to the expanded bed reactor as described in Example 1.
- the procedure was the same, and the catalyst and stock oil used in the comparative patent were the same as in Example 1.
- the operating conditions and test results in this comparative example are shown in Tables 3 and 4, respectively.
- Example 2 and Example 3 compared with Comparative Example 2, the contents of impurities 8 and N in the product are significantly decreased, and the recovery rates of gasoline and diesel having higher added values are also different.
- the degree of improvement the use of the process of the present invention to simultaneously use a fluidized bed containing an inner circulation zone for the hydrotreating of inferior feedstock oil feedstock can significantly improve product quality and light oil yield, and can provide a qualified catalytic cracking feedstock.
- the catalyst addition mode of the invention can ensure that the main reactor bubbling bed reactor has maintained a stable operating state, ensuring smooth operation of the device and stable product quality. 1
- Table 2 by using the bubbling bed reactor in which the internal circulation zone is increased as described in Example 1, the hydrodesulfurization, the hydrodenitrogenation effect, and the conversion rate of the residue can be effectively improved.
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Description
一种沸腾床反应器及其加氢方法 技术领域
本发明涉及一种反应器及其使用工艺， 具体地， 涉及一种沸腾床反应器及 其加氢方法。 背景技术
随着重质原油的大力开发和世界范围内石油产品需求结构的变化， 市场对 轻质燃料油的需求持续快速增长和对重质燃料油的需求迅速减少， 重油加工技 术已经成为炼油工业开发的重点。 重油加工技术主要包括脱碳和加氢两种工艺 过程：
脱碳工艺主要包括溶剂脱沥青、 焦化、 重油催化裂化等， 脱碳工艺虽然设 备投资低， 但液体产品收率低， 品质差， 不能满足目前的环保要求。 同时， 随 着原油重质化、 劣质化趋势越来越严重， 渣油收率占原油的比例逐年上升， 最 高可达 70wt%以上， 而最常使用的重、 渣油的脱碳工艺为焦化工艺， 该工艺副 产物是大量低附加值的焦炭。
加氢方法根据催化剂在反应器中存在状态的不同可具体分为固定床加氢、 移动床加氢、 悬浮床加氢和沸腾床加氢。 在包括加氢的工艺中， 由于使用高压 反应设备而投资高， 但产品质量好， 液体收率高， 可以实现重、 渣油的最大限 度轻质化。 目前比较成熟的渣油加氢工艺为固定床渣油加氢， 但该工艺受到原 料性质制约， 对原料的金属、 残炭等指标要求比较严格。 悬浮床和移动床技术 在重油加工过程中虽然都具有一定的优势， 但近年来发展较慢， 由于悬浮床加 氢尾油富集大量重金属， 给尾油的加工和利用带来很大困难； 移动床加氢技术 通常采用原料油与催化剂逆流或并流通过反应器的操作方式， 利用催化剂的初 活性进行重油加工， 虽然加氢效果较好， 但催化剂用量大， 并且催化剂的加氢 活性未被充分利用。
目前， 由于沸腾床加氢技术可以实现催化剂在线加排， 具有广泛的原料适 应性并能够保证长周期运行， 所以发展势头强劲。 沸腾床反应器是气、 液、 固 三相流化床， 可以处理高金属、 高沥青质含量的重、 劣质原料油， 其具有压力 降小， 温度分布均匀， 可保持整个运行周期内催化剂活性恒定， 可在运行过程
中加入新鲜催化剂和取出废催化剂等特点。
催化剂的在线加排技术是保证沸腾床产品质量， 操作条件平稳和运行周期 的关键技术。 目前沸腾床加氢技术的催化剂在线加入方式通常釆用气相输送、 液相输送方法或靠重力将固体催化剂从位于反应器上部的高压储罐直接加入到 沸腾床反应器。 而在沸腾床反应器中要保持催化剂处于良好的沸腾状态， 要求 反应器中的液体粘度、 反应压力、 气液流速和反应温度等保持恒定， 而直接将 新鲜催化剂加入到沸腾床反应器中容易引起上述条件的瞬间波动， 造成反应器 内的流体状态和操作条件的瞬间操作不平稳， 此外， 由于新鲜催化剂初活性很 高， 将其直接加入沸腾床反应器， 与劣质的重、 渣油原料接触混合， 会导致催 化剂迅速积碳， 活性损失快， 影响反应物流的加氢反应效果， 增加催化剂的置 换频率。
发明专利 CN101418222A、 CN1335357A、 CN101360808A均为劣质渣油处 理方法的现有技术， 其中 CN101418222A采用了沸腾床与悬浮床的联合反应装 置； CN1335357A采用了膨胀床与移动床的联合反应装置； CN101360808A包括 了至少两个串联的上流式反应器， 但这些现有技术均没有公开在反应器内的催 化剂不能达到要求的活性时， 催化剂的在线处理方法。
US4398852介绍了一种沸腾床反应器的催化剂在线加入方法。 该方法为首 先将催化剂加入到耐高压的盛催化剂的容器中， 然后将其充氢至反应压力， 随 后将位于催化剂容器与反应器连接管线上的阀门打开， 使催化剂依靠重力进入 沸腾床反应器。 该过程为靠重力直接将催化剂加入到沸腾床反应器， 该催化剂 的加入方式将使初活性的催化剂在与劣质原料接触时迅速积碳， 加速了失活速 率， 增加了催化剂的置换频率， 同时， 由于催化剂和氢气的预热温度低于反应 温度， 所以会造成沸腾床反应温度波动， 操作条件不平稳， 影响产品质量。
US Re 25770和 US4398852描述了典型的沸腾床工艺，在沸腾床反应器中设 置了内循环杯进行气液分离， 提高了液体的转化率。 但该工艺在实际应用中存 在以下不足： 反应器内催化剂藏量较少， 反应器空间利用率低； 循环油泵维护 保养费用较高， 而且一旦循环油泵工作失常或损坏， 会造成催化剂下沉聚集， 装置被迫停止运行； 反应器内液体产品在非催化加氢条件下停留时间过长， 在 高温下很容易进行二次热裂解反应结焦而降低产品质量。
专利 CN02109404.7和 CN101376092A分别介绍了一种新型的沸腾床反应 器， 采用三相分离器以及带有导向口的三相分离器进行气、 液、 固有效分离。 与典型的沸腾床反应器相比， 其具有结构简单、 操作容易和反应器利用率高等
特点。但由于沸腾床反应器高径比较大， 通常为 1： 6至 1： 8， 并且反应器内除 了位于顶部的三相分离器外， 绝大部分有效反应空间为空筒结构， 缺少强制传 质构造， 所以气、 液、 固之间的传质效果较差， 液相产品的加氢效果不显著， 产品质量较差。 此外， 沸腾床反应器为返混反应器， 部分未来得及反应的原料 将随着反应后的物流排出反应器， 使原料转化率相对较低。
现有技术中有一种含有两个以上反应段的沸腾床， 在沸腾床反应器中设置 有三相分离部件进行气、 液、 固分离， 同时实现加氢脱金属、 加氢脱硫、 加氢 脱氮多个反应的顺序排列， 其中每个反应段使用一至两种催化剂。 所述三相分 离部件由导流构件和挡流构件构成， 所述导流构件为两头开口的上窄下宽或上 宽下窄的锥筒或锥斗， 有上位和下位之分， 下位导流构件的上端和上位导流构 件的下端同轴相套。 此反应器结构实际为将两个反应器合并而成的一个大型反 应器， 去掉了反应器之间的管线和其它如分离器， 沉降罐等设备， 其优点是可 以合理利用热能， 但是还存在以下不足： 反应器体积庞大， 增加了运输、 安装、 日常操作和维护的难度； 在一个反应器中随着反应段的增加， 三相分离器数目 随之增加， 使得整个反应器结构复杂， 并且三相分离器越多占据的反应器空间 越大， 气、 液、 固有效的反应空间越小， 而高温高压的沸腾床反应器成本很高， 这将造成资源的不合理利用； 采用多段串联操作方式， 对操作和各段之间的气 液分布盘设计要求很高， 要求操作条件非常平稳， 如果装置出现瞬间波动， 将 影响三相分离器的分离效果， 则气液夹带的固体可堵塞分布盘， 影响装置的正 常平稳运行。 发明内容
针对上述现有技术的不足， 本发明提供了一种沸腾床反应器。 该沸腾床反 应器内增设若干个内循环区， 可以有效提高原料转化率。
为实现上述目的， 本发明釆用的技术方案是：
一种沸腾床反应器， 包括垂直于地面的反应器壳体、 位于壳体内上部的相 分离器， 在所述相分离器的下方设置有内循环区， 所述内循环区包括圆筒、 锥 形扩散段和导流结构， 所述圆筒和设置于圆筒下端的锥形扩散段位于反应器壳 体中， 所述导流结构位于锥形扩散段下端的反应器壳体内壁上， 所述的导流结 构为设置于反应器内壁的环形凸起结构， 其沿反应器轴线的纵切面为梯形、 弓 形、 半圆形、 三角形中的一种， 或其它可起到导向作用的等同替换、 改进的形 状， 均应包含在本发明的保护范围之内。 所述圆筒的上端为略向外扩散的喇叭
口结构。
本发明新型沸腾床反应器的具体结构为：
在所述反应器壳体底部设有原料入口和气液分布板。 在反应器壳体顶部设 有气体出口， 上部壳壁设有液体排出口， 所述液体排出口介于相分离器内筒上 端开口和外筒下端开口之间的反应器壳壁上， 用于将反应生成的气体和液体导 出。 所述的相分离器设置于壳体内上部空间内， 包括内径不同的两个同心圆筒： 内筒和外筒。 所述的内筒和外筒的上下两端全部开口， 外筒的上端开口高于内 筒的上端开口， 而外筒的下端开口也应高于内筒的下端开口。 所述内筒的下端 为一锥形扩散段， 该扩散段的开口 （即内筒的下端开口） 直径小于反应器的内 径； 所述外筒的下端同样为一锥形扩散段， 该扩散段的开口 （即外筒的下端开 口） 直径也小于反应器的内径。
所述相分离器的内筒构成分离器的中心管， 内筒与外筒之间的环状空间组 成相分离器的折流筒， 外筒与反应器内壁之间的环状空间为该相分离器的澄清 液体产品收集区， 所述内筒的下端开口为物流导入口， 内筒的下端开口与反应 器内壁构成的圆环状开口为相分离器的催化剂下料口， 分离出的固体微粒催化 剂从此处重新返回到催化剂床层。
所述的相分离器各组成部件的具体尺寸及相对位置， 均可以由本领域设计 人员根据所使用的催化剂尺寸、 反应器处理量、 反应条件及分离效果等具体要 求通过计算或者简单的试验予以确定， 或者可以采用本领域中公开的常规技术， 例如可以参照本申请人之前申请的专利 CN02109404.7或 CN101376092A。
所述的内循环区包括圆筒、 锥形扩散段和相邻的导流结构。 圆筒与锥形扩 散段相连接， 该锥形扩散段的下端开口直径小于反应器的内径， 紧邻圆筒锥形 扩散段的为导流结构， 三者结合构成一个内循环区。 根据反应器的高径比和转 化深度的要求可以在反应器中设置一个或多个内循环区， 优选设置 2~3个内循 环区， 其中不同的内循环区的圆筒内径可以相同或不同。 所述导流结构为设置 于反应器壳体内壁的环形凸起结构， 其沿反应器轴线的纵切面为梯形、 弓形、 半圆形、 三角形， 或其它形状可起到导向作用的等同替换、 改进等， 均应包含 在本发明的保护范围之内。
靠近相分离器一侧的导流结构与反应器壁交点处的切线与反应器内壁形成 的夹角称为覆盖角， 该覆盖角为锐角， 最好小于 60度； 与之相对的一侧， 即远 离相分离器的导流结构与反应器壁交点处的切线与反应器内壁形成的夹角称为 摩擦角， 该摩擦角亦为锐角， 最好小于 60度。 所述导流结构围成的导流口的直
径介于相分离器的外筒和内筒直径之间。
根据本发明的沸腾床加氢反应器， 在紧邻相分离器的下方也可以设置导流 结构， 所述导流结构位于反应器内中上部， 位于相分离器与内循环区之间。 所 述导流结构与内循环区中的导流结构相类似。
所述的气体排出口一般位于在反应器顶部中心处。
为了将分离出的澄清液体排出反应器， 液体排出口一般设置在反应器壳壁 的上部， 其位置应介于相分离器外筒的上端开口与下端开口之间。
所述相分离器的上部通常设有一定的缓冲空间， 相分离后的气体产物于此 富集并从气体排出口排出反应器。
一般说来， 反应器的径高比范围在 0.01〜0.1之间。
本发明的沸腾床反应器通常还包括至少一个从所述反应器排出催化剂的部 件， 和至少一个向所述反应器补充新鲜催化剂的部件。 所述补充新鲜催化剂的 部件通常设置于所述反应器的顶部位置， 而所述排出催化剂的部件通常位于所 述反应器底部附近。 例如在反应器壳体顶部设置催化剂添加管， 而在底部设置 催化剂排出管。 所述的催化剂置换系统及使用方法， 可以是任何适用的设备或 方法， 例如可参照美国专利 US3398085或 US4398852所述的方法。
为了使反应原料在反应器中与催化剂均匀接触， 一般还应在所述圆筒型反 应器壳体内的底部设置气液分布板， 气液分布板可以选用任何可以使气体和液 相均匀分布的结构， 例如可采用泡帽结构。
沸腾床反应器的内循环区的作用原理为： 物流通过不同的反应器横截面积， 流速发生变化。 沸腾床反应器内的物流由气、 液、 固三相构成， 即固态的催化 剂， 液态的反应物流， 气态的氢气及生成的轻质烃。 当反应器内的物流通过的 横截面积发生变化时， 气体和液体流速随之发生变化， 则靠气液夹带的催化剂 将发生快速提升或沉降； 液相原料在沸腾床反应器中随着反应的进行生成部分 轻组分， 该组分会随着氢气一直向上通过反应器， 而部分反应后的液相产物及 未反应的原料将与催化剂的运动状态相类似， 在反应器内横截面较小的流体加 速区快速向上流动， 在横截面积瞬间扩大处将发生与主物流流动方向相反的逆 流现象。
采用带有内循环区和相分离器的沸腾床反应器用于加氢处理劣质原料油， 可以提高液相重组分的转化率， 该沸腾床反应器的结构可以提高反应器内物流 之间的传质和传热效果。
本发明的另一个目的是提供一种沸腾床反应器的加氢方法， · 该工艺既能保
证沸腾床装置在催化剂加入时平稳运行， 从而保证装置的运行周期， 也能对沸 腾床反应后物流进行进一步加工， 从而提高产品质量。
本发明沸腾床反应器加氢方法的技术方案如下：
劣质原料油与氢气的混合物经加热炉加热后， 以向上流动的方式进入沸腾 床反应器进行催化加氢反应， 所得流出物经过气液分离后的部分液相产物进入 膨胀床反应器进一步反应， 所述膨胀床反应器通过管线与沸腾床反应器相连； 当所述沸腾床反应器中的催化剂达不到催化活性要求， 不能保证产品质量时， 需要补充新鲜的催化剂， 所需的催化剂由膨胀床反应器予以补充， 而膨胀床反 应器缺少的催化剂由新鲜催化剂添加罐予以补充。
本发明方法中所述的膨胀床反应器的床层膨胀率为 5v%~25v%， 优选为 10v%~25v%, 最优选为 15v%〜20v%。本发明中所述的膨胀率是指催化剂膨胀后 的料面高度与催化剂自然装填后的料面高度差同催化剂自然装填后料面的高度 的比值。 所述膨胀床反应器的操作条件为： 反应压力为 6〜30MPa， 优选为 10~18MPa; 反应温度为 350〜500°C， 优选为 380〜430°C； 空速为 O.l^h'1, 优选 为 Ι ^·1; 氢油体积比为 400〜2000， 优选为 600〜1500。
膨胀床反应器每次催化剂的加入量为每次沸腾床反应器催化剂需要在线加 入量的 2〜20倍， 当膨胀床反应器中的催化剂剩余量为单次沸腾床反应器在线置 换量的 0〜5倍时， 从设置在该反应器上部的催化剂添加罐进行补充。
所述气液分离后的液相产物进入膨胀床反应器的比例占反应后液相产物总 量的 5wt%〜70wt%， 较好为 10wt%〜50wt%。
本发明方法中所述的劣质原料油包括常压渣油、 减压渣油、 脱沥青油、 油 砂沥青、 稠原油、 煤焦油及煤液化重油等劣质原料中的一种或几种。
所述沸腾床反应器的操作条件为：反应压力为 6〜30MPa，优选为 10〜18MPa; 反应温度为 350〜500°C , 优选为 400〜450°C ; 空速为 0.1〜5^， 优选为 0.5~3h ; 氢油体积比为 400〜2000， 优选为 600~1500。
根据本发明提供的劣质原料油加氢方法， 上述反应器中使用的催化剂可以 为本领域的常规沸腾床加氢催化剂， 典型的催化剂的性质为： 以耐熔无机氧化 物为载体， 以第 VIB 族和 /或第珊族金属为活性组分， 催化剂的颗粒直径为 0.8mm, 颗粒长度为 3〜5mm， 基本物化性质与传统的固定床加氢催化剂基本相 同。本发明优选使用具有以下性质的催化剂： 催化剂颗粒直径为 0.1~0.8mm，优 选为 0.1〜0.4mm，催化剂含有 IB族和第環族活性加氢金属组分。载体为 A1203， 含有至少一种助剂， 助剂选自如下几种元素： B、 Ca、 F、 Mg、 P、 Si、 Ti等，
助剂含量为 0,5wt%〜5.0wt%。 催化剂的孔容为 0.6〜1.2mL/g， 孔径 <8nm的孔容 小于 0.03mL/g,—般为 0.005〜0.02 mL/g，平均孔径为 15〜30nm，孔径在 15~30nm 之间的孔容占总孔容的 50%以上， 一般为 50%〜70%。 比表面为 100〜300m2/g, 优选为 120〜240m2/g。
催化剂含 VIB族金属氧化物 （如 Mo03) 1.0wt%~20.0wt%, 优选为 3.0wt% 〜15.0wt%， 含第珊族金属氧化物 （如 MO 或 CoO) 0.1wt%~8.0wt%, 优选为 0.5wt%~5.0wt%。 催化剂磨耗≤0.1 wt%。
沸腾床反应器中所使用的催化剂为粒径在 0.1〜0.8mm之间的微球催化剂， 而目前常规的沸腾床反应器，如 US Re 25570及相关专利中公开的反应器中设置 的内循环杯主要功能是进行气液的有效分离， 反应器中使用的催化剂与常规的 加氢催化剂颗粒大小基本相同， 所以常规的沸腾床反应器不适用本发明方法。
本发明方法中使用的沸腾床反应器可以选择内部设置有气、 液、 固三相分 离器、及导流口等内构件的沸腾床反应器。如 CN1448212A、 CN101376092A中 公开的沸腾床反应器都可用于本发明的加氢方法。 上述的反应器虽然可以用于 本发明方法， 但由于沸腾床反应器的绝大部分有效反应区为空筒结构， 气液固 之间的传质效果差， 所以加氢效果不显著， 同时该返混操作的反应器会使部分 未来得及反应的进料将随着反应后物流流出反应器， 原料转化率相对较低。
本发明的劣质原料油加氢处理方法采用带有内循环区和三相分离器反应器 的沸腾床与膨胀床联合工艺处理重油原料， 既可以提高轻质油品质量， 也可以 保证主反应器沸腾床反应器在催化剂添加过程中平稳操作。 该组合工艺可以采 用灵活的操作模式进行重油加工， 首先重油原料经沸腾床加氢反应器进行加氢 裂化反应， 反应后物流进入分离装置得到气相产物和液相产物， 其中液相产物 的一部分循环至膨胀床反应器进行进一步加氢处理。 当沸腾床反应器需要进行 催化剂添加时， 膨胀床反应器物流夹带催化剂从反应器下部进入沸腾床反应器。 该灵活的操作过程克服了现有催化剂直接加入沸腾床反应器可能带来的反应器 温度和压力的波动问题， 消除了现有加剂方法对催化剂沸腾状态和反应物流性 质的影响 （例如造成操作不稳定， 如催化剂夹带或催化剂床层膨胀不理想， 影 响产品质量和装置的运行周期等）。而本发明将新鲜催化剂先加入到膨胀床反应 器， 然后再加入沸腾床反应器， 可以起到缓冲作用， 同时物流在膨胀床反应器 可以将催化剂预热至反应温度， 使得进入后续沸腾床反应器的液体、 催化剂和 气体的温度基本等于沸腾床的反应温度， 保证沸腾床装置的平稳运行。 此外， 首先使新鲜催化剂与沸腾床加氢后性质大为改进的物流接触， 可以充分利用催
化剂的初活性， 避免催化剂初期积碳过多而影响催化剂的效果。 本发明的有益效果：
本发明工艺简单、 科学、 合理， 与现有技术相比， 本发明的一种沸腾床反 应器及其加氢方法的优点为：
1) 本发明方法中提供的沸腾床反应器中形成一个或多个循环区， 可以形成 多个沸腾的操作区间， 使得整个沸腾床反应器操作更灵活。
2) 循环区间的存在延长了反应后物流中液体组分在沸腾床反应器中的停 留时间， 有利于增加轻质油收率。
3) 由于沸腾床反应器为返混程度很高的反应器， 所以通常反应器流出物中 含有部分未转化的原料， 设置多个小的内循环区， 可以使原料多次循环转化， 有利于提高转化率。
4) 沸腾床反应器的相分离器釆用圆筒形简单结构， 与现有技术中锥筒或锥 斗形相分离器相比， 具有制备工艺简单、 生产成本低、 便于安装和维修的优点。
5) 采用膨胀床反应器与沸腾床反应器联合的工艺， 与单一反应器相比较： 延长了反应路径和时间， 提高了反应物的杂质脱除率， 提高了产品质量。
6) 将原料的沸腾床加氢处理、 沸腾床反应后的物流在膨胀床进一步加氢及 催化剂的在线加入三种操作模式进行合理匹配， 既可以提高最终产品的质量， 也可以在充分发挥沸腾床加氢技术原料适应性广的前提下， 保证装置的长周期 平稳运行。
7) 新鲜催化剂先进入膨胀床反应器， 然后再进入沸腾床反应器， 可以起到 缓冲和预热催化剂的作用， 使得整个运行周期操作平稳， 同时使得新鲜催化剂 与性质大大改善的沸腾床加氢液体产物接触， 可以充分利用催化剂的初活性， 并能保证催化剂活性可稳定发挥。 附图说明
附图用来提供对本发明的进一步理解， 并且构成说明书的一部分， 与本发 明的实施例一起用于解释本发明， 并不构成对本发明的限制。 在附图中- 图 1为本发明沸腾床反应器的结构示意图 （只设置一个内循环区）。
图 2为本发明沸腾床反应器加氢方法的流程示意图。 结合附图， 本发明实施例中附图标记如下:
1一催化剂， 2—催化剂储罐， 3—膨胀床反应器， 4、 5、 11—氢气， 6—重 烃原料， 7—沸腾床反应器， 8—高压分离装置， 9一冷却净化装置， 10—蒸馏装 置， 12—汽油， 13—柴油， 14一加氢尾油， 15—催化剂排放管线， 16、 17、 18 一阀门；
101—进料口， 102—气液分布器， 103—反应器壳体， 104—导流结构， 105 一内筒， 106—外筒， 107—催化剂床层， 108—导流口， 109—催化剂进入口， 110—气体出口， 111一相分离器， 112—液体排出口， 113—下料口， 114—圆筒， 115—锥形扩散段， 116—催化剂排放口。 具体实施方式
以下结合附图对本发明的优选实施例进行说明， 应当理解为， 此处所描述 的优选实施例仅用于说明和解释本发明， 并不用于限定本发明。
如图 1 所示， 在本发明的一种具体实施方式中， 所使用的沸腾床反应器包 括内循环区和三相分离器， 其结构特征和工作原理为：
反应原料混合后由进料口 101进入反应器， 经气液分布器 102后均匀地通 过催化剂床层 107， 反应器壳体 103内的催化剂装量至少为反应器体积的 35%， 通常为 40%〜70%， 优选为 50%〜60%。在气液物流的携带作用下， 催化剂床层膨 胀到一定的高度， 其膨胀后体积通常比其静态体积大 20%〜70%。 进入反应区的 气液物流再与催化剂进行接触反应， 反应后的气液物流与未转化的原料及氢气 沿着反应器的轴线携带着固体催化剂上升进入由导流结构 104、圆筒 114及锥形 扩散段 115构成的循环区， 物流经导流结构 104形成的导流口 108和圆筒的锥 形扩散段 115汇集到圆筒 114流体通道中， 由于此时流体通道的横截面积减小， 所以气液流速加快， 流体通过圆筒顶部后， 流体通道瞬间扩大， 所以气液流体 流速瞬间减小， 其携带固体催化剂的能力降低， 致使部分反应后的液体和未转 化原料及固体催化剂沿着圆筒外壁与反应器内壁形成的通道向下流进入导流口 与从反应器下部向上流的物流混合， 从而形成小的循环区。 由循环区上升的气 相物流和部分液相物流及携带的催化剂进入紧邻相分离器的导流结构 104 围成 的导流口 108进入相分离器 111， 进行相分离： 气体首先分离出来， 通过气体排 出口 110排出反应器， 分离下来的催化剂经下料口 113返回反应区， 而基本不 含催化剂颗粒的澄清液相产物通过液体排出口 112 排出反应器。 为了及时将失 活的催化剂排出反应器和补充新鲜催化剂， 可以通过反应器上部的催化剂加入 管 109往反应系统中补充新鲜催化剂， 而通过反应器下部的排放管 116将部分
失活催化剂排出反应系统。
导流结构 104沿反应器轴线的纵切面为梯形， 覆盖角和摩擦角均为锐角， 最好均小于 60度。 当然导流结构 104沿反应器轴线的纵切面也可以为弓形或者 其他适宜的形状。
相分离器 111 由内径不同的同心圆筒内筒 105、 外筒 106连同反应器壳体 103的内壁共同构成。 内筒 105构成该相分离器的中心管， 内筒 105与外筒 106 之间的环状空间组成该相分离器的折流筒， 外筒 106与反应器壳体 103内壁之 间的环状空间为澄清液体产品收集区， 上述中心管下端扩散段的开口为物流导 入口， 该扩散段的开口与反应器壳体 103内壁构成的环状开口为催化剂下料口。 为了使折流筒内流体流速加快， 改善分离效果， 外筒扩散段锥形顶角一般比内 筒扩散段锥形顶角至少小 20度， 最好是小 40〜80度。
如图 2所示， 本发明的沸腾床反应器加氢方法过程为： 含重烃的原料 6和 氢气 5的混合物经加热炉加热后， 以上流式进入沸腾床反应器 7与催化剂接触 反应。 沸腾床加氢反应后物流从反应器顶部排出进入高压分离装置 8进行气液 分离， 部分分离出的液相产物与氢气 4混合以上流式进入膨胀床反应器 3进行 进一步加氢处理反应，反应产物从膨胀床反应器 3上部排出进入高压分离装置 8。 高压分离装置 8分离出的气体物流经冷却净化装置 9处理后， 气相可以作为循 环氢 11使用， 冷凝下来的轻组分与分离装置的部分液相物流混合进入蒸馏装置 10得到汽油 12, 柴油 13和加氢尾油 14。 加氢尾油可以作为催化裂化或固定床 渣油加氢原料， 或循环回沸腾床反应器 7。 当沸腾床反应器内催化剂活性降低， 不能提供要求质量的产品时， 需要进行催化剂置换， 此时的操作过程为： 将部 分沸腾床反应器中的失活催化剂经催化剂排放管线 15排出反应器， 然后开启位 于膨胀床反应器 3和沸腾床反应器 7连接管线上的阀门 18， 同时关闭位于膨胀 床反应器物料排出管线上的阀门 17， 使得含有固体催化剂的物流进入沸腾床反 应器 7， 在线加入催化剂的时间为 10〜50分钟。 沸腾床反应器 7催化剂添加结 束后， 恢复正常流程。 膨胀床反应器 3添加催化剂的方法为 ·· 首先将催化剂 1 装填到催化剂储罐 2中， 将该储罐充氢至略高于膨胀床压力 l〜5Pa， 打开催化 剂储罐 2和膨胀床 3之间的阀 16， 将新鲜催化剂加入到膨胀床反应器 3中。
为进一步说明本发明的技术方案和效果， 列举以下实施例。 其中涉及的百 分比均为重量百分比。 实施例 1
试验使用的渣油原料性质列于表 1。 由表 1 可知该渣油原料残炭值为 13.6wt%, 金属含量为 141.9μ ^， 沥青质为 6.4wt%， 硫含量为 2.5wt%， 氮含 量为 0.6wt%。
试验过程使用的微球形催化剂性质为： 平均直径为 0.6mm， 以氧化铝为载 体， 催化剂含 Mo03为 11.2wt%， 含 NiO为 3.0wt%， 含 P为 1.4wt%， 孔容为 0.60mL/g, 比表面积为 140m2/g， <8nm孔容占总孔容的 2.6%， 15〜30nm孔容占 总孔容的 65%。
采用常规的沸腾床加氢方法流程， 使用图 1 所示的带有一个内循环区的沸 腾床反应器。
沸腾床反应器内设有一个循环区， 沸腾床反应器的尺寸为： 反应器壳体的 内径 =160mm， 反应器壳体的高度 =3000mm， 壳体有效容积 60L， 分离器高度 =380mm, 分离器中心管圆柱部分直径 =92mm， 内筒下部锥形开口的底部直径 =144mm， 内筒下部锥体部分的高度 =41mm， 外筒圆柱部分直径 =128mm， 其锥 形部分开口的底部直径 =138mm，锥形部分的高度 =64mm， 外筒上部开口高于内 筒上部开口， 外筒下部锥形开口的底部位置高于内筒下部锥形开口的底部位置， 两者的高度差 =38mm， 分离器外筒上部开口与反应器壳体顶部切线的垂直距离 是 200mm， 液体产品管中心距反应器顶部切线的垂直距离是 338mm。环形导流 结构的覆盖角为 20°，摩擦角为 28°，紧邻相分离器的导向口直径为 100 mm。构 成循环区的导向口直径为 100mm, 圆筒内径为 80mm， 圆筒高度为 100mm， 锥 形扩散段的底部直径为 150mm， 锥形扩散段的高度为 45mm。 比较例 1
该比较例 1中反应器的基本结构同实施例 1，不同之处在于反应器内未设置 内循环区。 反应条件和试验原料同实施例 1， 其中具体试验条件及结果见表 2。 表 1 原料性质
项目 数据
密度 (20°C)， kg.m-3 1007.8
残炭值， wt% 13.6
粘度 (100°C)， mm2/s 576.7
凝点， 40
元素分析， wt%
C/ H 86.1/10.3
S/ N 2.5/0.6
金属元素， g.g-1
Fe/ Ni/ V 2.9/38.6/100.4
四组分分析， wt%
饱和烃 29.0
芳香烃 33.1
胶 质 31.5
沥青质 6.4 试验条件及试验结果
本实施例为本发明沸腾床反应器的一种加氢方法的实施例， 操作流程示意 图参照图 2， 沸腾床反应器内设有一个内循环区。
工艺过程为： 含重烃的原料 6和氢气 5混合后， 以上流式进入沸腾床反应 器 7与催化剂接触反应。 沸腾床加氢反应后物流从反应器顶部排出进入高压分 离装置 8进行气液分离， 将占反应后液相物流 15wt%的液相产物与氢气 4混合 以上流式进入膨胀床反应器 3进行进一步加氢处理反应， 反应产物从反应器上 部排出进入高压分离装置 8。高压分离装置分离出的气体物流经冷却净化装置 9
处理后， 气相可以作为循环氢 11使用， 冷凝下来的轻组分与分离装置的部分液 相物流混合进入蒸馏装置 10得到汽油 12， 柴油 13和加氢尾油 14。 当沸腾床反 应器内催化剂活性降低， 不能提供要求质量的产品时， 需要进行催化剂置换， 此操作过程为： 将部分沸腾床反应器中的失活催化剂经催化剂排放管线 15排出 反应器，然后开启位于膨胀床反应器 3和沸腾床反应器 7连接管线上的阀门 18， 同时关闭位于膨胀床反应器物料排出管线上的阀门 17， 使得含有固体催化剂的 物流进入沸腾床反应器 7， 在线加入催化剂的时间为 20分钟。 膨胀床反应器中 的催化剂添加方法为： 首先将催化剂 1装填到催化剂储罐 2中， 将该储罐充氢 至比膨胀床压力大约略高 2Pa， 打开催化剂储罐 2和膨胀床 3之间的阀 16， 将 新鲜催化剂加入到膨胀床反应器中。 其中膨胀床反应器中的催化剂床层膨胀高 度为 20v%， 膨胀床反应器内每次催化剂的加入量为每次沸腾床催化剂需要在线 加入量的 10倍， 当膨胀床反应器中的催化剂剩余量为单次沸腾床反应器在线置 换量的 4倍时， 从催化剂储罐 2加入新鲜催化剂。
沸腾床及膨胀床反应器的操作条件见表 3， 反应结果列于见表 4。 实施例 3
该实施例的工艺流程同实施例 2, 沸腾床反应器的基本结构同实施例 1， 只 是反应器内设置有两个循环区。 比较例 2
本实施例为本发明的一个对比实施方案， 其工艺过程与实施例 2基本相同， 不同之处在于未设置膨胀床， 当沸腾床反应器中的催化剂活性降低到不能提供 指标要求的产品而需要添加新鲜催化剂时， 直接从设置在沸腾床反应器顶部的 催化剂储罐向沸腾床反应器加入新鲜催化剂， 其加入过程与实施例 1 中所述的 由催化剂储罐向膨胀床反应器添加新鲜催化剂的过程相同， 此外， 该对比专利 中使用的催化剂和原料油与实施例 1 相同。 该比较例中的操作条件和试验结果 分别见表 3和 4。
反应条件
编 号 实施例 2 实施例 3 比较例 2 膨胀床反应器
反应温度， °c 425 422
反应压力， MPa 15 15
反应空速， h"1 1.0 1.0
氢油体积比 1500 1500
沸腾床反应器
反应温度， °c 425 422 425 反应压力， MPa 15 15 15 氢油体积比 700 700 700 反应空速， h"1 1.5 1.5 1.5 表 4 反应结果
最后应说明的是： 以上所述仅为本发明的优选实施例， 并不用于限制本发 明， 尽管参照前述实施例对本发明进行了详细的说明， 对于本领域的技术人员 来说， 其依然可以对前述各实施例所记载的技术方案进行修改， 或者对其中部 分技术特征进行等同替换。 凡在本发明的精神和原则之内， 所作的任何修改、 等同替换、 改进等， 均应包含在本发明的保护范围之内。
Claims
1. 一种沸腾床反应器， 包括垂直于地面的反应器壳体、 位于壳体内上部的 相分离器， 其特征在于， 在所述相分离器的下方设置有内循环区， 所述内循环 区包括圆筒、 锥形扩散段和导流结构， 所述圆筒和设置于圆筒下端的锥形扩散 段位于反应器壳体中， 所述导流结构位于锥形扩散段下端的反应器壳体内壁上， 所述的导流结构为设置于反应器内壁的环形凸起结构。
2. 根据权利要求 1所述的沸腾床反应器， 其特征在于， 所述内循环区的个 数为 2〜3个。
3. 根据权利要求 1所述的沸腾床反应器， 其特征在于， 所述相分离器与内 循环区之间设置有导流结构， 所述导流结构为设置于反应器内壁的环形凸起结 构。
4. 根据权利要求 1或 3所述的沸腾床反应器， 其特征在于， 所述导流结构 沿反应器轴线的纵切面为梯形、 弓形、 三角形或半圆形中的一种。
5. 根据权利要求 1所述的沸腾床反应器， 其特征在于， 所述导流结构的覆 盖角为锐角， 导流结构的摩擦角为锐角。
6. 根据权利要求 5所述的沸腾床反应器， 其特征在于， 所述的覆盖角和摩 擦角均小于 60度。
7. 根据权利要求 1所述的沸腾床反应器， 其特征在于， 所述的相分离器包 括内径不同的两个同心圆筒： 内筒和外筒， 所述内筒和外筒的上下两端全部开 口， 所述外筒的上端开口高于内筒的上端开口， 而外筒的下端开口也应高于内 筒的下端开口， 所述内筒的下端为一锥形扩散段， 该扩散段的开口直径小于反 应器的内径； 所述外筒的下端同样为一锥形扩散段， 该扩散段的开口直径也应 小于反应器的内径。
8. 根据权利要求 7所述的沸腾床反应器， 其特征在于， 外筒扩散段锥形顶 角比内筒扩散段锥形顶角小 20度〜 80度。
9. 根据权利要求 7所述的沸腾床反应器， 其特征在于， 所述的导流结构围 成的导流口的直径介于相分离器的外筒和内筒直径之间。
10. 根据权利要求 1 所述的沸腾床反应器， 其特征在于， 所述反应器壳体 内底部设置有分布板。
11. 根据权利要求 1 所述的沸腾床反应器， 其特征在于， 所述沸腾床的液 体排出口设置在反应器壳壁的上部， 介于相分离器内筒上端开口和外筒下端开 口之间的反应器壳壁上。
12. 根据权利要求 1 所述的沸腾床反应器， 其特征在于， 所述圆筒的上端 为略向外扩散的喇叭口结构。
13. 权利要求 1-12中的任意一项所述的沸腾床反应器的加氢方法， 其特征 在于， 包括如下步骤： 劣质原料和氢气的混合物首先在沸腾床反应器内进行催 化加氢反应， 所得产物中经过气液分离的部分液相产物进入膨胀床反应器在膨 胀床反应器内进一步反应， 所述膨胀床反应器通过管线与沸腾床反应器相连； 所述反应器壳体中的催化剂达不到催化活性要求时， 所需的催化剂由膨胀床反 应器予以补充。
14. 按照权利要求 13所述的沸腾床反应器的加氢方法， 其特征在于， 所述 沸腾床反应器的操作条件为： 反应压力为 6~30MPa， 反应温度为 350〜500°C， 空速为 0.1〜511 氢油体积比为 400〜2000。
15. 按照权利要求 13所述的沸腾床反应器的加氢方法， 其特征在于， 所述 膨胀床反应器的操作条件为： 床层膨胀率为 5v%〜25v%, 反应压力为 6〜30MPa， 反应温度为 350〜500°C , 空速为 O.l Shf1, 氢油体积比为 400~2000。
16. 按照权利要求 13所述的沸腾床反应器的加氢方法， 其特征在于， 所述 的气液分离后的液相产物进入膨胀床反应器的比例占沸腾床反应器反应后液相 产物总量的 5wt%〜70wt%。
17. 按照权利要求 16所述的沸腾床反应器的加氢方法， 其特征在于， 所述 的气液分离后的液相产物进入膨胀床反应器的比例占沸腾床反应器反应后液相 产物总量的 10wt%~50wt%。
18. 按照权利要求 13所述的沸腾床反应器的加氢方法， 其特征在于， 所述 膨胀床反应器每次的催化剂加入量为每次沸腾床反应器催化剂需要在线加入量 的 2~20倍， 当膨胀床反应器中的催化剂剩余量为单次沸腾床反应器在线置换量 的 0〜5倍时， 缺少的催化剂由新鲜催化剂添加罐进行补充。
19. 按照权利要求 13所述的沸腾床反应器的加氢方法， 其特征在于， 所述 的催化剂以耐熔无机氧化物为载体， 以第 VIB族和 /或第珊族金属为活性组分。
20. 按照权利要求 19所述的沸腾床反应器的加氢方法， 其特征在于， 所述 的催化剂具有如下性质：催化剂颗粒直径为 0.1〜0.8mm，催化剂含有 VIB族和第 珊族活性加氢金属组分， 载体为 A1203 ; 含有至少一种助¾], 助剂选自如下几种 元素： B、 Ca、 F、 Mg、 P、 Si、 Ti， 助剂含量为 0.5wt%~5.0wt%， 催化剂孔容 为 0.6〜1.2mL/g，平均孔径为 15~30nm，孔径在 15~30nm之间的孔容占总孔容的 50°/。以上， 孔径 <8nm的孔容小于 0.03mL/g， 比表面为 100~300m2/g。
21. 按照权利要求 14所述的沸腾床反应器的加氢方法， 其特征在于， 所述 沸腾床反应器的操作条件为： 反应压力为 10〜18MPa， 反应温度为 400〜450°C， 空速为 Ο^^ΐι·1, 氢油体积比为 600〜1500。
22. 按照权利要求 15所述的沸腾床反应器的加氢方法， 其特征在于， 所述 膨胀床反应器的操作条件为：床层膨胀率为 10ν%〜25ν%，反应压力为 10〜18MPa， 反应温度为 380〜430°C， 空速为 1 411·1, 氢油体积比为 600〜1500。
23. 按照权利要求 20所述的沸腾床反应器的加氢方法， 其特征在于， 所述 VIB 族加氢活性金属组分为 Mo， 以金属氧化物 1^003计含量为 1.0 wt%〜20.0wt%， 所述第 VI族加氢活性金属组分为 Ni或 Co， 含量以 NiO或 CoO 计为 0.1wt%〜8.0wt%。
24. 按照权利要求 23所述的沸腾床反应器的加氢方法， 其特征在于， 以金 属氧化物 Mo03计含量为 3.0 wt%~15.0wt% , 含量以 NiO 或 CoO 计为 0.5wt%〜5.0wt%。
25. 按照权利要求 20所述的沸腾床反应器的加氢方法， 其特征在于， 所述 的催化剂孔径 <8nm的孔容为 0.005〜0.02mL/g，孔径在 15〜30nm之间的孔容占 总孔容的 50%以上， 并小于等于 70%， 催化剂比表面为 120〜240m2/g。
26. 按照权利要求 13所述的沸腾床反应器的加氢方法， 其特征在于， 所述 的劣质原料油选自常压渣油、 减压渣油、 脱沥青油、 油砂沥青、 稠原油、 煤焦 油及煤液化重油中的一种或几种。
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Families Citing this family (21)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| CN103769010B (zh) * | 2012-10-19 | 2016-01-20 | 中国石油化工股份有限公司 | 一种沸腾床反应器 |
| CN103846063A (zh) * | 2012-11-29 | 2014-06-11 | 神华集团有限责任公司 | 一种催化剂分散器及含该催化剂分散器的反应器 |
| CN104560157B (zh) * | 2013-10-22 | 2016-06-22 | 中国石油化工股份有限公司 | 一种渣油加氢方法 |
| CN104560158B (zh) * | 2013-10-22 | 2016-07-20 | 中国石油化工股份有限公司 | 一种渣油加氢方法 |
| US9494503B2 (en) | 2013-11-06 | 2016-11-15 | Aspect Imaging Ltd. | Inline rheology/viscosity, density, and flow rate measurement |
| EP3151955B1 (en) * | 2014-06-09 | 2020-06-17 | Hatch Ltd. | Plug flow reactor with internal recirculation fluidized bed |
| US11300531B2 (en) | 2014-06-25 | 2022-04-12 | Aspect Ai Ltd. | Accurate water cut measurement |
| US10670574B2 (en) * | 2015-01-19 | 2020-06-02 | Aspect International (2015) Private Limited | NMR-based systems for crude oil enhancement and methods thereof |
| CN106053299B (zh) | 2015-04-12 | 2020-10-30 | 艾斯拜克特Ai有限公司 | 非圆形横截面管道中的流体的nmr成像 |
| CN106324010A (zh) | 2015-07-02 | 2017-01-11 | 艾斯拜克特Ai有限公司 | 使用mr设备对在管道中流动的流体的分析 |
| CN105363392B (zh) * | 2015-11-30 | 2017-12-01 | 华电重工股份有限公司 | 一种内环流式沸腾床反应器 |
| CN105457461B (zh) * | 2015-12-29 | 2018-04-10 | 原初科技（北京）有限公司 | 一种二氧化碳吸收并矿物化装置及方法 |
| US10655996B2 (en) | 2016-04-12 | 2020-05-19 | Aspect Imaging Ltd. | System and method for measuring velocity profiles |
| RU2650154C1 (ru) | 2016-12-16 | 2018-04-09 | Общество с ограниченной ответственностью "Биологические Источники Энергии" (ООО "БиоИстЭн") | Устройство с псевдоожиженным фонтанирующим слоем кольцеобразной формы и способ его работы |
| RU2658417C1 (ru) * | 2017-06-01 | 2018-06-21 | Федеральное государственное бюджетное образовательное учреждение высшего образования "Башкирский государственный университет" | Способ повышения эффективности селективного гидрирования |
| CN108148621A (zh) * | 2018-02-02 | 2018-06-12 | 上海英保能源化工科技有限公司 | 沸腾床加氢反应器及沸腾床加氢方法 |
| CN108554330B (zh) * | 2018-05-25 | 2024-04-02 | 北京中科诚毅科技发展有限公司 | 一种加氢反应器内导流工艺及其设计方法和用途 |
| US10682620B1 (en) | 2019-06-14 | 2020-06-16 | X Energy, Llc. | System for recovering entrained particles from an exhaust gas stream |
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Citations (15)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| USRE25570E (en) | 1964-05-05 | lemelson | ||
| USRE25770E (en) | 1965-04-27 | Gas-liquid contacting process | ||
| US3398085A (en) | 1965-09-14 | 1968-08-20 | Hydrocarbon Research Inc | Catalyst addition and withdrawal process |
| US4398852A (en) | 1980-04-21 | 1983-08-16 | Hri, Inc. | Process and apparatus for adding and removing particles from pressurized reactors |
| US5360535A (en) * | 1993-09-02 | 1994-11-01 | Texaco Inc. | Ebullated bed process with recycle eductor |
| CN2187492Y (zh) * | 1994-06-08 | 1995-01-18 | 天津大学 | 下喷自吸高效环流气液反应器 |
| CN1335357A (zh) | 2000-07-24 | 2002-02-13 | 中国石油化工股份有限公司 | 一种劣质重、渣油处理方法 |
| CN1448212A (zh) | 2002-04-04 | 2003-10-15 | 中国石油化工股份有限公司 | 一种沸腾床反应器 |
| CN101054534A (zh) * | 2006-03-08 | 2007-10-17 | Ifp公司 | 沸腾床中转化重油馏分及联产低硫中间馏分的方法和装置 |
| CN101240190A (zh) * | 2007-02-09 | 2008-08-13 | 中国石油化工股份有限公司 | 沸腾床加氢处理催化剂及其制备方法 |
| CN101356252A (zh) * | 2005-12-16 | 2009-01-28 | 雪佛龙美国公司 | 集成的重油改质方法和在线加氢精制方法 |
| CN101360808A (zh) | 2005-12-16 | 2009-02-04 | 雪佛龙美国公司 | 使用高活性浆料催化剂组合物进行重油改质的方法 |
| CN101376092A (zh) | 2007-08-27 | 2009-03-04 | 中国石油化工股份有限公司 | 一种新型沸腾床反应器 |
| CN101376834A (zh) * | 2007-08-27 | 2009-03-04 | 中国石油化工股份有限公司 | 一种沸腾床组合工艺 |
| CN101418222A (zh) | 2007-10-26 | 2009-04-29 | 中国石油化工股份有限公司 | 一种处理劣质渣油的组合工艺 |
Family Cites Families (5)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| US5925238A (en) * | 1997-05-09 | 1999-07-20 | Ifp North America | Catalytic multi-stage hydrodesulfurization of metals-containing petroleum residua with cascading of rejuvenated catalyst |
| CN1133720C (zh) | 2001-01-05 | 2004-01-07 | 中国石油化工股份有限公司 | 一种重、劣质原料油的加氢改质方法 |
| CN1205313C (zh) | 2002-05-15 | 2005-06-08 | 中国石油化工股份有限公司 | 一种渣油加工方法及设备 |
| US7938954B2 (en) * | 2005-12-16 | 2011-05-10 | Chevron U.S.A. Inc. | Systems and methods for producing a crude product |
| US7431831B2 (en) | 2005-12-16 | 2008-10-07 | Chevron U.S.A. Inc. | Integrated in-line pretreatment and heavy oil upgrading process |
-
2010
- 2010-10-19 CN CN201080047436.6A patent/CN102596386B/zh active Active
- 2010-10-19 WO PCT/CN2010/001641 patent/WO2011047538A1/zh active Application Filing
- 2010-10-19 US US13/502,218 patent/US9162207B2/en active Active
- 2010-10-19 EP EP10824379.1A patent/EP2492006A4/en active Pending
- 2010-10-19 RU RU2012119265/05A patent/RU2545330C2/ru active
- 2010-10-19 CA CA2778125A patent/CA2778125C/en active Active
Patent Citations (15)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| USRE25570E (en) | 1964-05-05 | lemelson | ||
| USRE25770E (en) | 1965-04-27 | Gas-liquid contacting process | ||
| US3398085A (en) | 1965-09-14 | 1968-08-20 | Hydrocarbon Research Inc | Catalyst addition and withdrawal process |
| US4398852A (en) | 1980-04-21 | 1983-08-16 | Hri, Inc. | Process and apparatus for adding and removing particles from pressurized reactors |
| US5360535A (en) * | 1993-09-02 | 1994-11-01 | Texaco Inc. | Ebullated bed process with recycle eductor |
| CN2187492Y (zh) * | 1994-06-08 | 1995-01-18 | 天津大学 | 下喷自吸高效环流气液反应器 |
| CN1335357A (zh) | 2000-07-24 | 2002-02-13 | 中国石油化工股份有限公司 | 一种劣质重、渣油处理方法 |
| CN1448212A (zh) | 2002-04-04 | 2003-10-15 | 中国石油化工股份有限公司 | 一种沸腾床反应器 |
| CN101356252A (zh) * | 2005-12-16 | 2009-01-28 | 雪佛龙美国公司 | 集成的重油改质方法和在线加氢精制方法 |
| CN101360808A (zh) | 2005-12-16 | 2009-02-04 | 雪佛龙美国公司 | 使用高活性浆料催化剂组合物进行重油改质的方法 |
| CN101054534A (zh) * | 2006-03-08 | 2007-10-17 | Ifp公司 | 沸腾床中转化重油馏分及联产低硫中间馏分的方法和装置 |
| CN101240190A (zh) * | 2007-02-09 | 2008-08-13 | 中国石油化工股份有限公司 | 沸腾床加氢处理催化剂及其制备方法 |
| CN101376092A (zh) | 2007-08-27 | 2009-03-04 | 中国石油化工股份有限公司 | 一种新型沸腾床反应器 |
| CN101376834A (zh) * | 2007-08-27 | 2009-03-04 | 中国石油化工股份有限公司 | 一种沸腾床组合工艺 |
| CN101418222A (zh) | 2007-10-26 | 2009-04-29 | 中国石油化工股份有限公司 | 一种处理劣质渣油的组合工艺 |
Non-Patent Citations (1)
| Title |
|---|
| See also references of EP2492006A4 |
Also Published As
| Publication number | Publication date |
|---|---|
| CA2778125A1 (en) | 2011-04-28 |
| CN102596386A (zh) | 2012-07-18 |
| CA2778125C (en) | 2018-10-02 |
| US9162207B2 (en) | 2015-10-20 |
| RU2545330C2 (ru) | 2015-03-27 |
| EP2492006A1 (en) | 2012-08-29 |
| US20120205288A1 (en) | 2012-08-16 |
| EP2492006A4 (en) | 2018-05-23 |
| CN102596386B (zh) | 2014-07-30 |
| RU2012119265A (ru) | 2013-11-27 |
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