CN111234869B - C for pyrolysis gasoline9+ fraction hydrogenation process - Google Patents

C for pyrolysis gasoline9+ fraction hydrogenation process Download PDF

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CN111234869B
CN111234869B CN201811448720.9A CN201811448720A CN111234869B CN 111234869 B CN111234869 B CN 111234869B CN 201811448720 A CN201811448720 A CN 201811448720A CN 111234869 B CN111234869 B CN 111234869B
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hydrogenation
reactor
catalyst
material flow
product
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CN111234869A (en
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杜周
熊凯
纪玉国
张富春
季静
李正艳
任玉梅
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Sinopec Beijing Research Institute of Chemical Industry
China Petroleum and Chemical Corp
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Sinopec Beijing Research Institute of Chemical Industry
China Petroleum and Chemical Corp
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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
    • C10G67/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only
    • C10G67/02—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only plural serial stages only
    • 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
    • C10G2300/20—Characteristics of the feedstock or the products
    • C10G2300/201—Impurities
    • C10G2300/202—Heteroatoms content, i.e. S, N, O, P

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  • Oil, Petroleum & Natural Gas (AREA)
  • Engineering & Computer Science (AREA)
  • Chemical Kinetics & Catalysis (AREA)
  • General Chemical & Material Sciences (AREA)
  • Organic Chemistry (AREA)
  • Production Of Liquid Hydrocarbon Mixture For Refining Petroleum (AREA)

Abstract

The invention discloses C for pyrolysis gasoline9 +The fraction hydrogenation method utilizes the hydrogenation process of sectional hydrogenation and circulating a part of the first product flow to the first reactor, and circulating a part of the second product flow to the colloid separation tower, the first reactor and the second reactor respectively, and is matched with a catalyst with high low-temperature activity, so that the operation stability of the device and the hydrogenation activity of the catalyst are improved. The method can reduce the export bromine number of the finally prepared hydrogenation product to be below 3g bromine/100 g oil, the export diene value to be 0g iodine/100 g oil and the sulfur content to be below 3 mg/kg.

Description

C for pyrolysis gasoline9+ fraction hydrogenation process
Technical Field
The invention relates to the field of pyrolysis gasoline hydrogenation, in particular to C for pyrolysis gasoline9 +A method for hydrogenating distillate.
Background
Pyrolysis gasoline is the most significant by-product of ethylene production, with the amount produced being about 50% of the ethylene yield. Pyrolysis gasoline is the major source of aromatics and has a composition that contains about 50% aromatics. C of pyrolysis gasoline9 +Fraction (hydrocarbons of carbon nine and above in pyrolysis gasoline, referred to as pyrolysis C for short)9 +) The mass fraction of (A) is about 10-20% of the ethylene yield. In recent years, with the rapid development of petrochemical industry in China, particularly the improvement of ethylene production capacity, cracking C9 +The yield of (2) is increasing. The full and reasonable utilization of the part of resources can have great influence on the improvement of the overall benefit of the ethylene device and the deep processing development of ethylene byproducts, and is also an important subject of the research of the ethylene post-processing industry at home and abroad at present. At present, only about 30% of cracking C9 +The obtained component is reasonably utilized as a synthetic resin, and the rest about 70 percent of the cracking C9 +Basically used as low-grade fuel, and has low utilization rate and additional value.
Cleavage C9 +The oil contains more than 200 kinds of hydrocarbons, has extremely complex composition, mainly comprises 50-60% of aromatic hydrocarbon and 25-40% of dicyclopentadiene by mass fraction, has the initial boiling point of 140-160 ℃, the final boiling point of 200-220 ℃, the density of 0.88-0.94 g/mL, the bromine number of less than or equal to 200g of bromine per 100g of oil, and contains higher impurities such as moisture, sulfur, oxygen, arsenic and the like and colloid with the concentration of more than 200mg/100 mL. Due to the lack of related technologies and the like, most of domestic ethylene plants will crack C9 +Sold as inexpensive primary raw materials or used as fuels, which are pre-processed by only a few downstream units as gasoline, diesel components or solvent oils. However, the catalysts used in these devices all present a counter cracking C9 +High raw material quality requirement, low space velocity, short one-way operation period and the like.
Cleavage C9 +The product obtained after hydrofining and separation can be used as a high-octane gasoline additive or a high-boiling-point aromatic hydrocarbon solvent oil and other products. However, with the increasing year by year of the processing amount of foreign crude oil in China, the quality of the crude oil is heavy and high-sulfuration, the yield of downstream by-products (pyrolysis gasoline) is increased gradually, and the quality is in a descending trend, which is mainly characterized in that the content of arsenic and sulfur impurities is greatly increased, and the components are heavy. Cleavage C9 +Belongs to heavy components in pyrolysis gasoline, and shows the characteristics of high bromine number, high colloid content and high sulfur content of raw materials, so the difficulty of hydrotreating is highIn C6～C8And (6) cutting. The conventional pyrolysis gasoline hydrogenation method cannot ensure pyrolysis C9 +The hydrogenation device can stably operate for a long time.
In the prior art, when selective hydrogenation treatment is carried out on full-fraction pyrolysis gasoline, the hydrogenation process condition is that the liquid volume space velocity is generally less than or equal to 4h-1The reactor inlet temperature is 40-130 ℃, the reaction pressure is more than or equal to 2MPa, the hydrogen-oil ratio is 100-500 (v/v), the catalyst is a nickel catalyst (prepared by taking alumina as a carrier and adopting an impregnation method, and the catalyst comprises 14-20% of nickel oxide, 1-8% of lanthanum oxide and/or cerium oxide and 1-8% of VIB oxide auxiliary agent, 2-8% of silicon dioxide, 1-8% of alkaline earth metal oxide and the specific surface area of the catalyst is 60m by taking the weight of the catalyst as 100%2/g～150m2(iii) a pore volume of 0.4mL/g to 0.6 mL/g). But because the method adopts one-stage hydrogenation treatment, the hydrogenation effect is poor, and the method is difficult to be applied to cracking C with high colloid content9 +In the hydrogenation process of (2).
Disclosure of Invention
In view of the problems of the prior art, it is an object of the present invention to provide a C for pyrolysis gasoline9 +The fraction hydrogenation method utilizes the hydrogenation process of sectional hydrogenation and circulating a part of the first product flow to the first reactor, and circulating a part of the second product flow to the colloid separation tower, the first reactor and the second reactor respectively, and is matched with a catalyst with high low-temperature activity, so that the operation stability of the device and the hydrogenation activity of the catalyst are improved.
In one aspect of the invention, C is provided for pyrolysis gasoline9 +A process for hydrogenating a distillate comprising:
providing C comprising pyrolysis gasoline9 +A feed stream to a distillate;
the raw material flow is subjected to degumming treatment in a colloid separation tower to obtain a degumming material flow;
the degumming material flow and hydrogen are subjected to a first-stage hydrogenation reaction under the action of a first catalyst in a first reactor to obtain a first product material flow;
a part of the first product stream flows into a second reactor as a first-stage hydrogenation product, and the rest part of the first product stream is recycled to the first reactor to form a first recycle stream;
the first-stage hydrogenation product flowing into the second reactor and hydrogen are subjected to second-stage hydrogenation reaction under the action of a second catalyst in the second reactor to obtain a second product material flow;
and collecting a part of the second product stream as a second-stage hydrogenation product, and respectively recycling the rest of the second product stream to the colloid separation tower to form a second recycle stream, to the first reactor to form a third recycle stream, and to the second reactor to form a fourth recycle stream.
The inventor of the application discovers in research that the colloid content of the raw material flow can be effectively reduced by carrying out ungluing treatment on the raw material flow, the occurrence of coking and carbon deposition of the catalyst in the subsequent hydrogenation process is reduced, the service life of the catalyst is effectively prolonged, and the operation time of the device is further prolonged. Moreover, by adopting sectional hydrogenation, in particular two-section hydrogenation process, the method can be used for the C of the pyrolysis gasoline9 +The hydrocarbons with higher unsaturated degree, which are easier to polymerize and hydrogenate, and the hydrocarbons with lower unsaturated degree, such as mono-olefin, which are less easy to polymerize and hydrogenate, in the fraction are subjected to hydrogenation treatment in sections, so that the hydrogenation efficiency can be improved, and the generation of byproducts can be reduced.
By carrying out the above-mentioned cycle,
in the colloid separation tower, a mixed material flow of the raw material flow and the second circulating material flow is formed, so that the colloid content in the mixed material flow entering the colloid separation tower can be effectively reduced, the difficulty of colloid separation can be further reduced, the colloid separation can be effectively realized at a lower temperature (on the premise of ensuring the colloid separation effect, the temperature of the top of the tower after circulation is established can be reduced by 15-25 ℃ compared with the temperature of the top of the colloid separation tower which is fed at the beginning), and the energy consumption can be reduced;
in the first reactor, a mixed material flow of the degelatinized material flow, the first circulating material flow and the third circulating material flow is formed, and the relationship of bromine number of the three components is that the degelatinized material flow is larger than the first circulating material flow and larger than the third circulating material flow, so that the bromine number of the mixed material flow existing in the first reactor can be reduced, and further, new colloid can be reduced and inhibited from being generated, and therefore, the coking and carbon deposition of the catalyst can be reduced and inhibited, the service life of the catalyst can be prolonged, and the long-term stable operation of a hydrogenation device can be ensured;
in the second reactor, a mixed material flow of the first-stage hydrogenation product and the fourth circulating material flow is formed, and the bromine number of the fourth circulating material flow is smaller than that of the first-stage hydrogenation product, so that the bromine number of the mixed material flow existing in the second reactor can be reduced, new colloid can be reduced and inhibited from being generated, coking and carbon deposition of the catalyst can be reduced and inhibited, the service life of the catalyst can be prolonged, and the long-term stable operation of a hydrogenation device can be ensured.
Preferably, the second recycle stream and the feed stream are mixed and then introduced into the colloid separation tower; mixing the degummed material flow, the first circulating material flow and the third circulating material flow and then introducing the mixture into a first reactor; and mixing the fourth circulating material flow and the first-stage hydrogenation product, and introducing into a second reactor.
In a preferred embodiment of the present invention, the volume ratio of the first recycle stream to the degummed stream is 1 (2-10), preferably 1 (3-8); the volume ratio of the second recycle material flow to the raw material flow is 1 (1-4), preferably 1 (1-2); the volume ratio of the third circulating material flow to the degumming material flow is 1 (1-4), and 1 (1-2) is preferred; the volume ratio of the fourth recycle stream to the first-stage hydrogenation product is 1 (1-4), and preferably 1 (1-2).
According to the invention, the above limitation is beneficial to both the efficiency and stability of colloid separation and the efficiency and stability of hydrogenation reaction.
According to the invention, the method further comprises cooling the first recycle stream, the second recycle stream, the third recycle stream and the fourth recycle stream before recycling the first recycle stream, the second recycle stream, the third recycle stream and the fourth recycle stream to the colloid separation tower, the first reactor and the second reactor respectively.
Preferably, the temperature of the cooled recycle stream is lower than the inlet temperature of the reactor. Since the hydrogenation reaction is an exothermic reaction, the temperature in the reactor increases with the progress of the hydrogenation reaction, which leads to polymerization of unsaturated hydrocarbons, coking and carbon deposition of the catalyst, and the like. The cooling treatment is carried out on the circulating material flow, so that the reaction temperature is favorably controlled, and the adverse effect caused by overhigh temperature of a catalyst bed layer is avoided.
The cooling process is not limited, and may be any of those methods for reducing the temperature of the recycle stream, which are used in the prior art, and in the present application, the distribution and cooling of the recycle stream are performed by a cooling drum.
In another preferred embodiment of the present invention, the temperature of the bottom of the colloid separation tower is 150 to 190 ℃, the temperature of the top of the colloid separation tower is 70 to 120 ℃, and the vacuum degree is 60 to 90 kPa.
In another preferred embodiment of the present invention, the first catalyst comprises Al supported on2O3/TiO2An active component on the composite oxide support, the active component being selected from at least one of nickel and nickel oxide; preferably, the content of the active component is 12 to 25 percent based on the total mass of the first catalyst; based on the Al2O3/TiO2Total mass of the composite oxide support, Al2O3The content of (A) is 70% -90%, preferably 75% -85%; TiO 22The content of (A) is 10% -30%, preferably 15% -25%.
According to the invention, a specific Al with good low-temperature activity and loaded with nickel and/or nickel oxide is used2O3/TiO2The composite oxide carrier is used as the first catalyst and is favorable for cracking C9 +The unsaturated hydrocarbons with high degree of unsaturation, such as methyl styrene, indene, dicyclo-cyclopentadiene, long-chain diolefin and the like which are easy to polymerize and easy to hydrogenate, reach a saturated state through the hydrogenation reaction, and effectively avoid the occurrence of polymerization reaction and the generation of undesired products such as colloid and the like.
In another preferred embodiment of the present invention, the first stepThe two catalysts comprise Al supported on the Al2O3/TiO2An active component on the composite oxide carrier, wherein the active component is selected from at least one of nickel, molybdenum, cobalt and oxides thereof; preferably, based on the total mass of the second catalyst, the content of nickel and/or nickel oxide is 0-8%, the content of molybdenum and/or molybdenum oxide is 8-20%, and the content of cobalt and/or cobalt oxide is 1-5%; based on the Al2O3/TiO2Al is based on the total weight of the composite oxide support2O3The content of (A) is 70% -90%, preferably 70% -80%; TiO 22The content of (A) is 10% -30%, preferably 20% -30%.
According to the present invention, preferably, the active components of the second catalyst may be molybdenum (or an oxide thereof) and cobalt (or an oxide thereof); nickel (or its oxide), molybdenum (or its oxide) and cobalt (or its oxide) are also possible.
According to the invention, by using specific Al loaded with active components such as nickel, molybdenum, cobalt and the like2O3/TiO2The composite oxide carrier is used as a second catalyst and is beneficial to cracking C9 +In (2), hydrocarbons having a relatively low degree of unsaturation, such as monoolefins, which are relatively difficult to polymerize and hydrogenate, are hydrogenated.
In another preferred embodiment of the present invention, in the one-stage hydrogenation reaction, the inlet temperature of the first reactor is 30 ℃ to 100 ℃, preferably 40 ℃ to 90 ℃; the reaction pressure is 2.0MPa-5.0MPa, preferably 2.6MPa-3.2 MPa; the feed space velocity is 0.5h-1-10h-1Preferably 0.5h-1-10h-1(ii) a The hydrogen-oil ratio is (100-.
According to the invention, the feed space velocity refers to the space velocity of the stream entering the reactor.
According to the invention, the hydrogen-to-oil ratio refers to the volume ratio of hydrogen to the stream entering the reactor.
In another preferred embodiment of the present invention, in the two-stage hydrogenation reaction, the inlet temperature of the second reactor is 200 ℃ to 300 ℃, preferablyIs 220-280 ℃; the reaction pressure is 2.0MPa-5.0MPa, preferably 2.6MPa-3.2 MPa; the feed space velocity is 0.5h-1-10h-1Preferably 0.5h-1-10h-1(ii) a The hydrogen-oil ratio is (300-.
According to the invention, hydrogenation of different components can be realized in stages by first carrying out a first-stage hydrogenation reaction at a lower temperature and then carrying out a second-stage hydrogenation reaction at a higher temperature. Specifically, at a lower temperature, hydrogenation reaction is carried out on hydrocarbons which are easy to polymerize and easy to hydrogenate and have a higher unsaturated degree, such as methyl styrene, indene, dicyclo-cyclopentadiene, long-chain diolefin and the like, and meanwhile, polymerization is not caused due to the lower temperature, and undesired side products such as colloid and the like are generated; at a relatively high temperature, the unsaturated hydrocarbon which is relatively easy to polymerize is hydrogenated to a saturated state so as not to be easily polymerized, only hydrogenation reaction is carried out, and the hydrocarbon which is relatively difficult to polymerize and not easy to hydrogenate, such as monoolefine and other hydrocarbons with lower unsaturation degree, is hydrogenated under relatively violent reaction conditions. This can promote the progress of the hydrogenation reaction while reducing or avoiding the occurrence of the polymerization reaction.
In another preferred embodiment of the present invention, further comprising,
a step of subjecting the first catalyst to a reduction treatment before the first-stage hydrogenation reaction; and
and a step of subjecting the second catalyst to a sulfidation treatment before the second-stage hydrogenation reaction.
According to the invention, the reduction treatment comprises: under the hydrogen pressure of 0.5MPa to 2.0MPa, the temperature is raised to 120 ℃ to 160 ℃ at the temperature raising rate of 30 DEG/h to 45 DEG/h, and after the temperature is maintained for 3h to 7h, the temperature is raised to 430 ℃ to 470 ℃ at the temperature raising rate of 20 DEG/h to 30 DEG/h, and the temperature is maintained for 10h to 20 h.
According to the invention, the vulcanization treatment comprises: under the hydrogen pressure of 2.0MPa-3.5MPa, the hydrogen-oil ratio of (300--1-5h-1The stream volume space velocity of (A) is introduced into a sulfurised oil containing dimethyl disulfide (DMDS) for 10h to 30 h.
According to the present invention, the hydrogen-oil ratio referred to in the vulcanization treatment means a volume ratio of hydrogen gas to the vulcanized oil.
According to the invention, the first catalyst is subjected to reduction treatment, so that the nickel oxide can be reduced into elemental nickel, and the hydrogenation activity of the first catalyst is improved. And the hydrogenation activity of the second catalyst can be improved by carrying out sulfurization treatment on the second catalyst.
In another preferred embodiment of the present invention, the first product stream has a bromine number of less than 40g bromine per 100g oil.
According to the present invention, by controlling the bromine number of the first product stream within the above range, it is possible to avoid that the one-stage hydrogenation reaction proceeds too vigorously to cause undesirable side reactions such as polymerization or to affect the stability of the hydrogenation apparatus.
In another preferred embodiment of the present invention, the specific surface area of the second catalyst is 200m2/g-300m2The pore volume is 0.5ml/g-0.7 ml/g.
According to the present invention, by limiting the specific surface area and pore volume of the second catalyst to the above ranges, the second-stage hydrogenation reaction can be more efficiently performed.
The hydrogenation method of the invention is suitable for the C aiming at the pyrolysis gasoline9 +The distillate is subjected to deep hydrotreatment. Moreover, by adopting the hydrogenation method, the export bromine number of the finally prepared hydrogenation product can be reduced to be below 3g bromine/100 g oil, the export diene value reaches 0g iodine/100 g oil, the sulfur content is reduced to be below 3mg/kg, and the requirement of the market on the quality of the hydrogenation product is met.
Drawings
FIG. 1 shows a schematic diagram of the hydrogenation apparatus and the flow direction of the streams used in example 1.
Reference numerals: i-colloid separation tower, II-first reactor, III-first cold separating tank, IV-second reactor, V-second cold separating tank, 1-raw material flow, 2-colloid removing flow, 3-first product flow, 4-first circulating flow, 5-first-stage hydrogenation product, 6-second product flow, 7-second-stage hydrogenation product, 8-fourth circulating flow, 9-third circulating flow, 10-second circulating flow and 11-fresh hydrogen.
Detailed Description
The present invention will be described in detail below with reference to examples, but the scope of the present invention is not limited to the following description.
The hydrogenation products obtained in examples 1 to 2 and comparative examples 1 to 2 were tested and analyzed according to the following methods or standards.
Bromine number: the test was performed according to GB/T1815-1997.
Sulfur content: the test was performed according to SH/T0253-1992.
Diene value: testing was performed according to ASTM UOP 326-2008.
Example 1
Step A, preparation of the first catalyst
a)Al2O3/TiO2Preparation of composite oxide support
The specific surface area is taken to be 160m2A pore volume of 0.58ml/g and a maximum pore diameter of
Figure BDA0001883568110000071
170g of clover alumina, putting the clover alumina into 106ml of cyclohexane solution of tetraethyl titanate (the mass concentration of the tetraethyl titanate is 0.808ml/g), stirring for 15 minutes, drying at 120 ℃ for 6 hours, and roasting at 600 ℃ for 4 hours to prepare Al2O3/TiO2A composite oxide support (designated Z1). Wherein, TiO2The content is 15.0%.
b) Loading of active ingredients
100g of Z1 obtained in step a) was taken and immersed in 50ml of an aqueous nickel nitrate solution (mass concentration of nickel nitrate was 24g/100ml in terms of nickel atom) for 2 hours, dehydrated with compressed air for 20 minutes, dried at 110 ℃ for 8 hours, and then calcined at 550 ℃ for 5 hours to obtain a first catalyst precursor (named NZ 1).
68g of the NZ1 was taken out and immersed in 50ml of an aqueous nickel nitrate solution (the mass concentration of nickel nitrate was 16g/100ml in terms of nickel atom) for 1 hour, dehydrated with compressed air for 20 minutes, dried at 110 ℃ for 8 hours, and then calcined at 550 ℃ for 5 hours, to obtain the first catalyst. Wherein the nickel content in the first catalyst is 25 wt% in terms of nickel atoms.
Step B, preparation of the second catalyst
a)Al2O3/TiO2Preparation of composite oxide support
201.3g of AlCl3·6H2Placing O (analytically pure) in 1000ml of deionized water to prepare solution A1; 421.7g of Ti (OCH)2CH3) (chemical purity) was dissolved in 500ml of benzene (benzene content 99.8 wt%) to obtain solution B1; 18g of NH4HCO3(analytically pure) is dissolved in 600ml of deionized water, then 250ml of ammonia water with the concentration of 26 wt% is added, the mixture is stirred evenly, and a certain volume of deionized water is added to prepare 1000ml of solution C1.
Under the condition of normal pressure and 73 ℃, three solutions of A1, B1 and C1 are subjected to co-current co-precipitation. Controlling the flow rate of the solution C1 to keep the pH value of the precipitate within the range of 5.0-6.0 for 8 minutes, increasing the flow rate of the solution C1 to keep the pH value of the mixed solution within the range of 8.5-9.5 for 8 minutes, then reducing the flow rate of the solution C1 to keep the pH value of the mixed solution within the range of 5.0-6.0 for 8 minutes, increasing the flow rate of the solution C1 to keep the pH value of the precipitate within the range of 8.5-9.5, and repeating the steps until all the solutions A1 and B1 are added dropwise to prepare the reaction solution.
The obtained reaction solution was allowed to stand at 70 ℃ for 30 minutes and then filtered to obtain a filter cake. Then washing the filter cake with deionized water 15 times the volume of the filter cake for 30 minutes, filtering again, washing again, repeating the process four times, finally drying the filter cake at 110 ℃ for 10 hours, and roasting at 450 ℃ for 5 hours to obtain 42.7g of Al2O3/TiO2A composite oxide support (designated Z2). Wherein, TiO2The content of (B) was 15.14%.
b) Loading of active ingredients
Preparing 26.68g/100ml of ammonium molybdate tetrahydrate by using a volumetric flask and 14 wt% of ammonia water, impregnating 100g of the carrier Z2 prepared in the step a) at normal temperature for 2 hours, filtering, drying at 110 ℃ overnight, and roasting at 550 ℃ for 4 hours to complete the loading of the active component molybdenum oxide.
Then putting the molybdenum oxide-loaded Z2 into 22.34g/100ml cobalt nitrate hexahydrate solution, soaking for 2h at normal temperature, filtering, drying overnight at 110 ℃, and roasting for 4h at 550 ℃ to obtain MoO3MoO with a content of 15 wt.% and a CoO content of 4 wt%3-CoO/Al2O3-TiO2The second catalyst of (1).
Step C, pretreatment of the catalyst
a) Hydroprocessing of a first catalyst
And (C) filling the first catalyst prepared in the step A in a first reactor, and carrying out hydrotreating on the first catalyst. The treatment conditions include: the temperature was raised to 140 ℃ under a hydrogen pressure of 0.5MPa at a temperature raising rate of 45 DEG/h and held at that temperature for 5 hours, and thereafter, the temperature was raised to 450 ℃ at a temperature raising rate of 30 DEG/h and held at that temperature for 15 hours.
b) Sulfidation treatment of the second catalyst
And (C) filling the second catalyst prepared in the step (B) in a second reactor and carrying out vulcanization treatment on the second catalyst. The treatment conditions include: under the hydrogen pressure of 2.0MPa, the hydrogen-oil ratio of 330:1 and the temperature of 350 ℃, for 2h-1Was charged with 20h of 4 wt% DMDS cyclohexane sulfide oil.
Step D, cracking C of gasoline9 +Hydrogenation of distillate
The material flow is cracking C produced by Dushan Tianli industries9 +The raw material, bromine number is 68g bromine/100 g oil, sulfur content is 223mg/kg, diene value is 6.2g iodine/100 g oil.
And (3) introducing a raw material flow (1) into the colloid separation tower (I) to carry out colloid removal treatment. The degelatinization treatment conditions include:
the bottom temperature of the colloid separation tower is 190 ℃, the top temperature of the colloid separation tower is 120 ℃, and the vacuum degree of the colloid separation tower is 80 kPa. Through colloid separation treatment, a degumming material flow (2) separated from the top of a colloid separation tower is mixed with fresh hydrogen (11) (the hydrogen-oil ratio is 200:1), then flows into the top inlet of a first reactor (II), passes through a filled first catalyst bed layer, and has the inlet temperature of 40 ℃, the reaction pressure of 2.8MPa and the feeding space velocity of 2.8MPa1.8h-1And a first product stream (3) is obtained from the bottom outlet of the first reactor.
The first product flow is cooled by a first cold separating tank (III) and is divided into two parts, one part of the first product flow is used as a first-stage hydrogenation product (5) to flow into the top inlet of the second reactor, and the rest part of the first product flow is used as a first circulating flow (4) and is recycled to the first reactor after being mixed with a degumming flow and fresh hydrogen, wherein the volume ratio of the first circulating flow to the degumming flow is 1: 4.
The first-stage hydrogenation product is mixed with fresh hydrogen (the hydrogen-oil ratio is 600:1), flows into the top inlet of a second reactor (IV), passes through a filled second catalyst bed layer, and has the inlet temperature of 220 ℃, the reaction pressure of 2.7MPa and the feeding space velocity of 1.8h-1And a second product stream (6) is obtained from the bottom outlet of the second reactor.
The second product stream was cooled and split into four parts by a second cold split pot (v), the first part was collected as the second stage hydrogenation product (7) and tested for bromine number, sulfur content and diene number at different time points, with the results shown in table 1; the second part is taken as a second recycle stream (10) and recycled to the colloid separation tower after being mixed with the feed stream, wherein the volume ratio of the second recycle stream to the feed stream is 1: 1; a third part is taken as a third recycle stream (9) and is recycled to the first reactor after being mixed with the degumming material stream, wherein the volume ratio of the third recycle stream to the degumming material stream is 1: 1; the remaining part is recycled to the second reactor as a fourth recycle stream (8) after being mixed with the first-stage hydrogenation product, wherein the volume ratio of the fourth recycle stream to the first-stage hydrogenation product is 1: 2. The bromine number (5d) of the first product stream was 20g of bromine per 100g of oil and the bromine number of the second product stream (second stage hydrogenation product) is shown in Table 1.
TABLE 1
Figure BDA0001883568110000101
In addition, after the circulation was established, the overhead temperature of the colloid separation column was adjusted to 100 ℃.
Example 2
Step A, preparation of the first catalyst
a)Al2O3/TiO2Preparation of composite oxide support
The specific surface area is taken to be 160m2A pore volume of 0.58ml/g and a maximum pore diameter of
Figure BDA0001883568110000102
170g of clover alumina, placing the clover alumina in 100ml of cyclohexane solution of tetraethyl titanate (the mass concentration of the tetraethyl titanate is 0.490ml/g), stirring for 15 minutes, drying at 120 ℃ for 6 hours, and roasting at 600 ℃ for 4 hours to prepare Al2O3/TiO2A composite oxide support (designated Z3). Wherein, TiO2The content was 24.3%.
b) Loading of active ingredients
100g of Z2 obtained in step a) was taken and immersed in 50ml of an aqueous nickel nitrate solution (mass concentration of nickel nitrate was 20g/100ml in terms of nickel atom) for 2 hours, dehydrated with compressed air for 20 minutes, dried at 110 ℃ for 8 hours, and then calcined at 550 ℃ for 5 hours to obtain a first catalyst precursor (named NZ 3).
73g of the prepared NZ2 was taken and placed in 50ml of a nickel nitrate aqueous solution (the mass concentration of nickel nitrate is 20g/100ml in terms of nickel atom) to be impregnated for 2 hours, and after dehydration treatment with compressed air for 20 minutes, the catalyst was dried at 110 ℃ for 8 hours and then calcined at 550 ℃ for 5 hours, thereby preparing the first catalyst. Wherein the nickel content in the first catalyst is 16 wt% in terms of nickel atom.
Step B, preparation of the second catalyst
a)Al2O3/TiO2Preparation of composite oxide support
178.5g of AlCl3·6H2Placing O (analytically pure) in 1000ml of deionized water to prepare solution A2; 430.7g of Ti (OCH)2CH3) (chemical purity) was dissolved in 500ml of benzene (benzene content 99.8 wt%) to obtain solution B1; 18g of NH4HCO3(analytically pure) deionised in 600mlAdding 250ml of ammonia water with the concentration of 26 wt% into water, uniformly stirring, and adding a certain volume of deionized water to prepare 1000ml of solution C1.
Under the condition of normal pressure and 73 ℃, three solutions of A1, B1 and C1 are subjected to co-current co-precipitation. Controlling the flow rate of the solution C1 to keep the pH value of the precipitate within the range of 5.0-6.0 for 8 minutes, increasing the flow rate of the solution C1 to keep the pH value of the mixed solution within the range of 8.5-9.5 for 8 minutes, then reducing the flow rate of the solution C1 to keep the pH value of the mixed solution within the range of 5.0-6.0 for 8 minutes, increasing the flow rate of the solution C1 to keep the pH value of the precipitate within the range of 8.5-9.5, and repeating the steps until all the solutions A1 and B1 are added dropwise to prepare the reaction solution.
The obtained reaction solution was allowed to stand at 70 ℃ for 30 minutes and then filtered to obtain a filter cake. Then washing the filter cake with deionized water 15 times the volume of the filter cake for 30 minutes, filtering again, washing again, repeating the process four times, finally drying the filter cake at 110 ℃ for 10 hours, and roasting at 450 ℃ for 5 hours to obtain 42.7g of Al2O3/TiO2A composite oxide support (designated Z4). Wherein, TiO2The content of (B) was 15.14%.
b) Loading of active ingredients
Preparing 26.68g/100ml of ammonium molybdate tetrahydrate by using a volumetric flask and 14 wt% of ammonia water, impregnating 100g of the carrier Z4 prepared in the step a) at normal temperature for 2 hours, filtering, drying at 110 ℃ overnight, and roasting at 550 ℃ for 4 hours to complete the loading of the active component molybdenum oxide.
Then putting the molybdenum oxide-loaded Z2 into 39.48g/100ml nickel nitrate hexahydrate and 11.62g/100ml cobalt nitrate hexahydrate solution, soaking for 2h at normal temperature, filtering, drying at 110 ℃ overnight, and roasting at 550 ℃ for 4h to obtain MoO3MoO with a content of 14.1 wt%, a CoO content of 1.9 wt% and a NiO content of 7.5%3-CoO/Al2O3-TiO2The second catalyst of (1).
Step C, pretreatment of the catalyst
a) Hydroprocessing of a first catalyst
And (C) filling the first catalyst prepared in the step A in a first reactor, and carrying out hydrotreating on the first catalyst. The treatment conditions include: the temperature was raised to 150 ℃ under a hydrogen pressure of 0.5MPa at a temperature raising rate of 40 DEG/h and held at that temperature for 5 hours, and thereafter, the temperature was raised to 480 ℃ at a temperature raising rate of 35 DEG/h and held at that temperature for 10 hours.
b) Sulfidation treatment of the second catalyst
And (C) filling the second catalyst prepared in the step (B) in a second reactor and carrying out vulcanization treatment on the second catalyst. The treatment conditions include: under the hydrogen pressure of 2.5MPa, the hydrogen-oil ratio of 400:1 and the temperature of 330 ℃, for 1h-1Was charged with 18h of 2.5 wt% DMDS cyclohexane sulfide oil.
Step D, cracking C of gasoline9 +Hydrogenation of distillate
The material flow is cracking C produced by Beijing Yanshan division9 +The raw material, bromine number is 81.28g bromine/100 g oil, sulfur content is 150mg/kg, diene value is 5.86g iodine/100 g oil.
And introducing the raw material flow into the colloid separation tower for degumming treatment. The degelatinization treatment conditions include:
the bottom temperature of the colloid separation tower is 170 ℃, the top temperature of the colloid separation tower is 100 ℃, and the vacuum degree of the colloid separation tower is 90 kPa. Through colloid separation treatment, a colloid-removed material flow separated from the top of a colloid separation tower is mixed with fresh hydrogen (the hydrogen-oil ratio is 300:1), flows into a top inlet of a first reactor, passes through a filled first catalyst bed layer, and has the inlet temperature of 32 ℃, the reaction pressure of 2.8MPa and the feeding airspeed of 1.5h-1Under the conditions of (1), carrying out a first-stage hydrogenation reaction, and obtaining a first product stream from a bottom outlet of the first reactor.
The first product flow is cooled by a first cold knockout drum and divided into two parts, one part of the first product flow is used as a first-stage hydrogenation product to flow into the top inlet of the second reactor, and the rest part of the first product flow is used as a first circulating flow to be recycled to the first reactor after being mixed with the degumming flow and fresh hydrogen, wherein the volume ratio of the first circulating flow to the degumming flow is 1: 2.
The first-stage hydrogenation product is mixed with fresh hydrogen (the hydrogen-oil ratio is 800:1) and flows into a second reactorThe top inlet of the reactor passes through the packed second catalyst bed layer, the inlet temperature is 210 ℃, the reaction pressure is 2.7MPa, and the feeding space velocity is 1.5h-1And carrying out a second-stage hydrogenation reaction under the conditions of (1) and obtaining a second product material flow from a bottom outlet of the second reactor.
The second product stream was cooled in a second cold trap and divided into four fractions, the first fraction was collected as the second-stage hydrogenation product and tested for bromine number, sulfur content and diene number at different time points, the results are shown in table 2; the second part is taken as a second recycle stream and recycled to the colloid separation tower after being mixed with the raw material stream, wherein the volume ratio of the second recycle stream to the raw material stream is 1: 2; the third part is taken as a third recycle material flow and is recycled to the first reactor after being mixed with the degumming material flow, wherein the volume ratio of the third recycle material flow to the degumming material flow is 1: 2; the rest part is taken as a fourth recycle stream and recycled to the second reactor after being mixed with the first-stage hydrogenation product, wherein the volume ratio of the fourth recycle stream to the first-stage hydrogenation product is 1: 1. The bromine number (5d) of the first product stream was 20g of bromine per 100g of oil and the bromine number of the second product stream (second stage hydrogenation product) is shown in Table 2.
TABLE 2
Figure BDA0001883568110000131
In addition, after the circulation was established, the overhead temperature of the colloid separation column was adjusted to 80 ℃.
Comparative example 1
The reaction conditions were the same as in example 1, except that HTC200 was used as the first catalyst, LD145+ HR306C was used as the second catalyst, and the entire second product stream was used as the second-stage hydrogenation product (i.e., recycle of stream was only performed in the first-stage hydrogenation reaction, and not in the second-stage hydrogenation reaction). The hydrogenation products were tested for bromine number, sulfur content and diene number at various time points and the results are shown in table 3.
TABLE 3
Figure BDA0001883568110000132
In addition, after the circulation was established, the overhead temperature of the colloid separation column was adjusted to 100 ℃.
Comparative example 2
Hydrogenation was carried out using the catalyst of example 2 and in the manner of example 2, except that the entire second product stream was taken as the second-stage hydrogenation product (i.e. recycling of the stream was carried out only in the first-stage hydrogenation reaction, and not in the second-stage hydrogenation reaction). The hydrogenation products were tested for bromine number, sulfur content and diene number at various time points and the results are shown in table 4.
TABLE 4
Figure BDA0001883568110000141
In addition, after the circulation was established, the overhead temperature of the colloid separation column was adjusted to 100 ℃.
As can be seen by comparing the data in examples 1-2 and comparative examples 1-2 described above, the present invention provides C for pyrolysis gasoline9 +The hydrogenation method of the distillate can still obtain hydrogenation products meeting the market requirements (the market requirements are that bromine number is less than or equal to 3g bromine/100 g oil, sulfur content is less than or equal to 3mg/kg, and diene value is 0g iodine/100 g oil) after the device operates for 90 days, and has better stability. The hydrogenation method of comparative example 1 cannot obtain a hydrogenation product satisfying the market requirements; the hydrogenation process of comparative example 2, after 35 days of operation of the apparatus, did not produce a hydrogenation product meeting the market requirements.
It should be noted that the above-mentioned embodiments are only for explaining the present invention, and do not set any limit to the present invention. The present invention has been described with reference to exemplary embodiments, but the words which have been used herein are words of description and illustration, rather than words of limitation. The invention can be modified, as prescribed, within the scope of the claims and without departing from the scope and spirit of the invention. Although the invention has been described herein with reference to particular means, materials and embodiments, the invention is not intended to be limited to the particulars disclosed herein, but rather extends to all other methods and applications having the same functionality.

Claims (18)

1. C for pyrolysis gasoline9 +A process for hydrogenating a distillate comprising:
providing C comprising pyrolysis gasoline9 +A feed stream to a distillate;
the raw material flow is subjected to degumming treatment in a colloid separation tower to obtain a degumming material flow;
the degumming material flow and hydrogen are subjected to a first-stage hydrogenation reaction under the action of a first catalyst in a first reactor to obtain a first product material flow;
a part of the first product stream flows into a second reactor as a first-stage hydrogenation product, and the rest part of the first product stream is recycled to the first reactor to form a first recycle stream;
the first-stage hydrogenation product flowing into the second reactor and hydrogen are subjected to second-stage hydrogenation reaction under the action of a second catalyst in the second reactor to obtain a second product material flow;
collecting a part of the second product stream as a second-stage hydrogenation product, and respectively recycling the rest of the second product stream to the colloid separation tower to form a second recycle stream, to the first reactor to form a third recycle stream, and to the second reactor to form a fourth recycle stream,
wherein the volume ratio of the first recycle stream to the degummed stream is 1 (2-10); the volume ratio of the second circulating material flow to the raw material flow is 1 (1-4); the volume ratio of the third circulating material flow to the degumming material flow is 1 (1-4); the volume ratio of the fourth recycle stream to the first-stage hydrogenation product is 1 (1-4).
2. The hydrogenation process of claim 1, wherein the volume ratio of the first recycle stream to the degummed stream is 1 (3-8); the volume ratio of the second recycle material flow to the raw material flow is 1 (1-2); the volume ratio of the third circulating material flow to the degumming material flow is 1 (1-2); the volume ratio of the fourth recycle stream to the first-stage hydrogenation product is 1 (1-2).
3. The hydrogenation method according to claim 1 or 2, wherein the bottom temperature of the colloid separation tower is 150-190 ℃, the top temperature of the colloid separation tower is 70-120 ℃, and the vacuum degree of the colloid separation tower is 60-90 kPa.
4. The hydrogenation process of claim 1 or 2, wherein the first catalyst comprises Al supported2O3/TiO2An active component on the composite oxide support, the active component being selected from at least one of nickel and nickel oxide.
5. The hydrogenation process according to claim 4, wherein the active component is present in an amount of 12 to 25% by mass, based on the total mass of the first catalyst.
6. The hydrogenation process of claim 4, wherein the Al is based on2O3/TiO2Total mass of the composite oxide support, Al2O3The content of (A) is 70-90%; TiO 22The content of (A) is 10% -30%.
7. The hydrogenation process of claim 4, wherein the Al is based on2O3/TiO2Total mass of the composite oxide support, Al2O3The content of (A) is 75-85%; TiO 22The content of (A) is 15% -25%.
8. The hydrogenation process of claim 1 or 2, wherein the second catalyst comprises Al supported2O3/TiO2An active component on the composite oxide carrier, wherein the active component is at least one of nickel, molybdenum, cobalt and oxides thereof.
9. Hydrogenation process according to claim 8, characterized in that the content of nickel and/or nickel oxides, the content of molybdenum and/or molybdenum oxides and the content of cobalt and/or cobalt oxides is between 0% and 8%, between 8% and 20% and between 1% and 5% based on the total mass of the second catalyst.
10. The hydrogenation process of claim 8, wherein Al is based on2O3/TiO2Al is based on the total weight of the composite oxide support2O3The content of (A) is 70-90%; TiO 22The content of (A) is 10% -30%.
11. The hydrogenation process of claim 8, wherein Al is based on2O3/TiO2Al is based on the total weight of the composite oxide support2O3The content of (A) is 70-80%; TiO 22The content of (A) is 20-30%.
12. The hydrogenation process of claim 1 or 2, wherein in the first stage hydrogenation reaction, the inlet temperature of the first reactor is from 30 ℃ to 100 ℃; the reaction pressure is 2.0MPa-5.0 MPa; the feed space velocity is 0.5h-1-10h-1(ii) a The hydrogen-oil ratio is (100- & ltSP & gt 600- & gt) 1.
13. The hydrogenation process of claim 12, wherein in the first stage hydrogenation reaction, the inlet temperature of the first reactor is from 40 ℃ to 90 ℃; the reaction pressure is 2.6MPa-3.2 MPa; the feed space velocity is 0.5h-1-1.8h-1(ii) a The hydrogen-oil ratio is (200- & gt 400) & gt 1.
14. The hydrogenation process of claim 1 or 2, wherein, in the second-stage hydrogenation reaction, the inlet temperature of the second reactor is from 200 ℃ to 300 ℃; the reaction pressure is 2.0MPa-5.0 MPa; the feed space velocity is 0.5h-1-10h-1(ii) a The hydrogen-oil ratio is (300-:1。
15. The hydrogenation process of claim 14, wherein, in the second-stage hydrogenation reaction, the inlet temperature of the second reactor is from 220 ℃ to 280 ℃; the reaction pressure is 2.6MPa-3.2 MPa; the feed space velocity is 0.5h-1-1.8h-1(ii) a The hydrogen-oil ratio is (400- & 800): 1.
16. The hydrogenation method according to claim 1 or 2, further comprising,
a step of subjecting the first catalyst to a reduction treatment before the first-stage hydrogenation reaction; and
and a step of subjecting the second catalyst to a sulfidation treatment before the second-stage hydrogenation reaction.
17. The hydrogenation process of claim 1 or 2, wherein the first product stream has a bromine number of less than 40g bromine per 100g of oil.
18. The hydrogenation process according to claim 1 or 2, wherein the second catalyst has a specific surface area of 200m2/g-300m2The pore volume is 0.5ml/g-0.7 ml/g.
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