Method for dechlorinating and decoloring petroleum resin

Abstract

The invention relates to the field of petroleum resin catalytic hydrogenation decolorization, and discloses a petroleum resin dechlorination decolorization method which comprises the following steps: (A) in the presence of a dechlorination catalytic adsorbent and in a protective atmosphere, performing a dechlorination reaction on a petroleum resin mixed solution to obtain a dechlorination product; (B) in the presence of a hydrogenation catalyst, mixing the dechlorination product with hydrogen for hydrogenation reaction to obtain a hydrogenation product; wherein the petroleum resin mixed liquid comprises a petroleum resin raw material and a diffusion liquid; the diffusion liquid is selected from at least one of C6-C12 monocyclic alkane, monocyclic aromatic hydrocarbon and C5-C10 paraffin. According to the method, the diffusion liquid is adopted for dechlorinating the petroleum resin under the protective atmosphere, so that the petroleum resin raw material is more convenient to convey and diffuse on the surface and inside the catalyst, the chlorine content in the petroleum resin is greatly reduced, the corrosion of chlorine to the device is prevented, the service life of the device is prolonged, and the operation risk is reduced.

Description

Technical Field

[0001] This invention relates to the field of catalytic hydrogenation decolorization of petroleum resins, and more specifically to a method for dechlorination and decolorization of petroleum resins. Background Technology

[0002] Petroleum resins, also known as hydrocarbon resins, are produced by processing C5 and C9 fractions from ethylene produced by petroleum cracking. After a series of processes including raw material pretreatment, polymerization, catalyst removal, solvent removal, and curing, petroleum resins can be obtained in solid or viscous liquid states. Petroleum resins are oligomer mixtures among thermoplastic resins, with molecular weights generally around 300-3000 and softening temperatures between 50-150℃. They are characterized by low acid value, low melting point, good miscibility, good adhesion, and resistance to water, acids, and alkalis. According to the different raw materials, petroleum resins can be roughly divided into five types: (1) aliphatic petroleum resins based on C5 or C9 fractions; (2) aromatic petroleum resins based on C9 fractions; (3) copolymer petroleum resins based on C5 and C9 fractions; (4) DCPD (dicyclopentadiene) alicyclic petroleum resins; and (5) petroleum resins modified by hydrogenation.

[0003] Petroleum resin molecules contain numerous unsaturated bonds, making them easily oxidized and exhibiting yellow, dark yellow, or even yellowish-brown colors, with unsatisfactory stability. Halogenated catalysts (such as aluminum trichloride) are frequently used in petroleum resin synthesis, resulting in the presence of halogens, metals, and other heteroatoms in the product. In particular, the presence of double bonds, benzene rings, and trace amounts of sulfur and chlorine impurities affects the resin's color, stability, and compatibility with other resins, limiting its applications. Hydrogenation is a crucial method for resin modification. Catalytic hydrogenation can saturate unsaturated bonds, remove sulfur, halogens, and other heteroatoms, thereby reducing the color, acid value, and bromine value of the petroleum resin, resulting in a near-colorless hydrogenated petroleum resin with excellent performance.

[0004] In the removal of halogens, especially chlorine, hydroremoval is the most commonly used method. While hydroremoval can remove chlorine relatively thoroughly, the generated hydrogen chloride is highly corrosive to the equipment. Particularly during the polymerization of C5 petroleum resin, a large amount of aluminum trichloride is used as a catalyst, introducing a significant amount of chlorine. Therefore, during hydrorefining and decolorization, a large amount of HCl gas is released. Simultaneously, petroleum resin contains a small amount of oxygen and is easily oxidized, resulting in an even higher oxygen content. This oxygen forms water after hydrogenation. When the dew point temperature of water is reached in the hydrogenation reactor, clear water is produced, which rapidly dissolves the large amount of HCl generated during hydroremoval, forming hydrochloric acid, a strong acid that severely corrodes the steel materials used in the equipment. Furthermore, petroleum resin hydrogenation is carried out under high temperature and high pressure conditions, and the reactors are all made of steel. Currently, no steel can withstand more than 10 ppm of HCl in a hydrogenation environment with water.

[0005] In hydrogenation, fixed-bed hydrogenation is highly popular due to its simple operation, high degree of continuity, and low cost. Currently, domestically, reduced-state metal catalysts are mainly used. Petroleum resins contain certain amounts of sulfur and chlorine impurities, which poison both noble metal catalysts and reduced-state nickel-based catalysts. Therefore, most hydrogenation processes employ two-stage or three-stage hydrogenation: first, a protective catalyst is used to remove sulfur and chlorine impurities through hydrogenation, followed by decolorization using a reduced-state metal catalyst. However, this results in the aforementioned severe corrosion problems in the upstream hydrogenation unit.

[0006] CN102746458A discloses a method for preparing hydrogenated C9 petroleum resin, employing a two-stage hydrogenation process. The first stage uses macroporous zeolite or diatomaceous earth as a carrier, with sulfided cobalt and molybdenum as active components. This method exhibits good resistance to poisoning and impurities, excellent stability, and a long service life. However, its drawbacks include the corrosion of the equipment caused by HCl dissolving in the generated water to form hydrochloric acid. Furthermore, the reaction temperature is relatively high, making the resin prone to decomposition and compromising product quality.

[0007] CN111548450A discloses a hydrogenation method for high-chlorine C9 petroleum resin, employing a three-stage hydrogenation process. The first stage involves hydrogenation for dechlorination and adsorption of HCl. The second stage involves hydrogenation for desulfurization and adsorption of sulfur. The third stage involves hydrogenation for decolorization. This patent uses hydrogenation adsorption for both dechlorination and desulfurization. While this can meet temporary dechlorination and desulfurization needs, the chlorine content in resins polymerized using AlCl3 as a catalyst is typically 1000-5000 mg / kg. If adsorption is used for removal, the adsorbent quickly reaches saturation, requiring replacement. Since this occurs in a hydrogen atmosphere and shares a hydrogen system with subsequent hydrogenation units, shutdown is necessary during adsorbent replacement. This significantly reduces the operational cycle of the industrial plant due to the adsorbent's chlorine penetration capacity.

[0008] CN115246912A discloses a method for producing low-color C9 petroleum resin. First, bleaching clay is used to remove the gum from the resin. Then, in a first-stage hydrogenation process, sulfur and chlorine are removed by hydrogenation using a sulfide-state catalyst, followed by adsorption with an adsorbent. Next, it enters a second-stage hydrogenation process using a metallic catalyst for decolorization. Finally, it enters a third-stage hydrogenation process for further purification. The method states that "the resin liquid after the first-stage hydrogenation is fed into an adsorption desulfurization and dechlorination tower filled with adsorption desulfurizing and dechlorination agents to remove hydrogen sulfide and hydrogen chloride from the resin liquid." However, it does not clearly specify whether the subsequent adsorption tank / tower operates under the same or different conditions as the first-stage hydrogenation. If they are under the same conditions, similar to CN111548450A, the same problem exists: once the adsorbent is saturated, the process must be stopped and replaced. The pressure of the first-stage hydrogenation is 10-16 MPa. Under high-pressure hydrogen conditions, it is impossible to switch adsorption tanks to extend the processing cycle. Switching tanks / reactors via valves under high-pressure hydrogen conditions poses significant safety risks. Under different operating conditions, in the first-stage hydrogenation reactor, where chlorine can be effectively removed under hydrogenation conditions, oxygen will also be removed, generating HCl and water respectively. If, under reactor operating conditions, these substances are not immediately adsorbed and removed after their formation, but instead flow out of the reactor with the feedstock and enter the next adsorption stage, significant hydrochloric acid corrosion will occur in the pipelines and equipment along the way, severely impacting equipment safety and production cycles. Furthermore, the second- and third-stage hydrogenation processes use metallic catalysts, and their lifespan and activity are entirely influenced by the removal efficiency of sulfur and chlorine in the first-stage hydrogenation, thus placing even higher demands on the adsorbent.

[0009] Therefore, finding a suitable dechlorination method is crucial. Summary of the Invention

[0010] The purpose of this invention is to overcome the problems of conventional petroleum resin dechlorination and decolorization methods, which use hydroremoval, generate HCl under high temperature and pressure, causing severe corrosion to the reactor, and the high risk of switching between reactors under high temperature and pressure after the dechlorination adsorbent becomes saturated. This invention provides a method for petroleum resin dechlorination and decolorization. This method uses a diffusion liquid under a protective atmosphere to dechlorinate the petroleum resin, making the transfer of petroleum resin raw materials and diffusion on and inside the catalyst surface more convenient, significantly reducing the chlorine content in the petroleum resin, preventing chlorine corrosion of the equipment, extending the service life of the equipment, and reducing operational risks.

[0011] To achieve the above objectives, the present invention provides a method for dechlorination and decolorization of petroleum resin, wherein the method includes the following steps:

[0012] (A) In the presence of a dechlorination catalytic adsorbent and under a protective atmosphere, a petroleum resin mixture is subjected to a dechlorination reaction to obtain a dechlorination product;

[0013] (B) In the presence of a hydrogenation catalyst, the dechlorination product is mixed with hydrogen gas to carry out a hydrogenation reaction to obtain the hydrogenated product;

[0014] The petroleum resin mixture includes petroleum resin raw materials and a diffusion liquid;

[0015] The diffusion liquid is selected from at least one of C6-C12 monocyclic alkanes, monocyclic aromatics, and C5-C10 alkanes.

[0016] Preferably, the conditions for the dechlorination reaction include: a reaction pressure of 1-3 MPa, more preferably 1.5-2.5 MPa; a reaction temperature of 150-250°C, more preferably 180-230°C; and a volume hourly space velocity of 1-6 h⁻¹. -1 Preferably 4-5 hours -1 .

[0017] Preferably, the conditions for the hydrogenation reaction include: a reaction temperature of 200-270℃, more preferably 200-240℃; a reaction pressure of 3-15 MPa, more preferably 5-13 MPa; and a volume hourly space velocity of 1-3 h⁻¹. -1 Preferably 1-2.5h -1 .

[0018] Preferably, the conditions for the hydrogenation reaction further include a hydrogen-to-oil volume ratio of 50-200, more preferably 80-180.

[0019] The beneficial effects achieved through the above technical solution are as follows:

[0020] (1) This invention provides a method for dechlorination and decolorization of petroleum resin. This method uses a diffusion liquid to dechlorinate petroleum resin under a protective atmosphere, which makes the internal and external diffusion of petroleum resin raw materials in the catalyst more convenient, and can significantly reduce the chlorine content in petroleum resin, prevent chlorine from corroding the hydrogenation unit under high temperature and high pressure, extend the service life of the unit, and reduce the operational risk.

[0021] (2) In this invention, preferably, the feed for the hydrogenation reaction in the method adopts an upflow mode, that is, the reactants enter from the bottom and exit from the top. Compared with the traditional trickle bed reactor with the inlet and outlet from the bottom, when the pressure drop occurs in the upflow reaction, the effect of the pressure drop can be eliminated by backflushing the gas path, that is, blowing gas from the top to the bottom of the reactor, without the need to stop the operation, which can extend the cycle of the project. The upflow reactor requires less hydrogen, which can significantly reduce the hydrogen-to-oil ratio, reduce the amount of hydrogen circulating, reduce the load on the circulating hydrogen compressor, and achieve the purpose of energy saving and consumption reduction. In addition, in the hydrogenation reaction, the sulfides in the raw materials are hydrogenated and desulfurized to generate hydrogen sulfide, which has an inhibitory effect on the activity of the catalyst. By adopting the upflow feed mode, the hydrogen sulfide gas generated in the reactor can more easily escape from the bottom to the top of the reaction liquid zone in the reactor, which greatly reduces the inhibition of the hydrogenation catalyst activity by hydrogen sulfide.

[0022] (3) In this invention, preferably, the hydrogenation catalyst, after activation, has good toxicity resistance and long lifespan, and compared with precious metal catalysts, the hydrogenation catalyst of this invention is a non-precious metal catalyst with low cost.

[0023] (4) In this invention, preferably, the method is combined with a dechlorination catalyst and a hydrogenation catalyst to process high-chlorinated petroleum resin, achieve dechlorination of petroleum resin at low temperature and low pressure, realize online switching of multiple dechlorination reactors, and have high reaction safety. After hydrogenation, the Gardner color of petroleum resin is reduced to less than 1. While achieving the purpose of dechlorination and decolorization, the softening point of petroleum resin does not change much. Attached Figure Description

[0024] Figure 1 This describes the activation status of the hydrogenation catalyst in Example 1 of the present invention;

[0025] Figure 2 This describes the activation status of the hydrogenation catalyst in Example 2 of the present invention;

[0026] Figure 3 This describes the activation status of the hydrogenation catalyst in Example 3 of the present invention. Detailed Implementation

[0027] The endpoints and any values ​​of the ranges disclosed herein are not limited to the precise ranges or values, and these ranges or values ​​should be understood to include values ​​close to these ranges or values. For numerical ranges, the endpoint values ​​of the various ranges, the endpoint values ​​of the various ranges and individual point values, and individual point values ​​can be combined with each other to obtain one or more new numerical ranges, which should be considered as specifically disclosed herein.

[0028] This invention provides a method for dechlorinating and decolorizing petroleum resin, wherein the method includes the following steps:

[0029] (A) In the presence of a dechlorination catalytic adsorbent and under a protective atmosphere, a petroleum resin mixture is subjected to a dechlorination reaction to obtain a dechlorination product;

[0030] (B) In the presence of a hydrogenation catalyst, the dechlorination product is mixed with hydrogen gas to carry out a hydrogenation reaction to obtain the hydrogenated product;

[0031] The petroleum resin mixture includes petroleum resin raw materials and a diffusion liquid;

[0032] The diffusion liquid is selected from at least one of C6-C12 monocyclic alkanes, monocyclic aromatics, and C5-C10 alkanes.

[0033] In this invention, the method significantly reduces the chlorine content in petroleum resin by dechlorinating it under a protective atmosphere, preventing chlorine from corroding the equipment. Combined with the subsequent hydrogenation reaction process, the upflow feeding mode can solve the problem of pressure drop during the operation of the hydrogenation reactor, which forces it to shut down, thus reducing operational risks.

[0034] According to the present invention, preferably, the petroleum resin raw material in the petroleum resin mixture has a mass percentage content of 10-40 wt%, more preferably 25-35 wt%; and the diffusion liquid has a mass percentage content of 60-90 wt%, more preferably 65-75 wt%.

[0035] In this invention, the petroleum resin has the conventional meaning in the art, and is preferably selected from at least one of C5 petroleum resin, C9 petroleum resin and DCPD (dicyclopentadiene) petroleum resin.

[0036] In this invention, the source of the petroleum resin raw material is not particularly limited. Preferably, the petroleum resin raw material is a polymer produced by processing C5 and C9 fraction by-products from petroleum cracking to produce ethylene, and is preferably selected from at least one of aliphatic petroleum resins based on C5 or C9 fractions, aromatic petroleum resins based on C9 fractions, copolymeric petroleum resins based on C5 and C9 fractions, and DCPD alicyclic petroleum resins.

[0037] According to the present invention, preferably, the diffusion liquid is selected from at least two of C6-C12 monocyclic alkanes, monocyclic aromatics, and C5-C10 alkanes.

[0038] According to the present invention, preferably, the diffusion liquid contains 65-90 wt% C6-C12 monocyclic alkanes, more preferably 75-85 wt%; 0-0.5 wt% monocyclic aromatics, more preferably 0-0.3 wt%; and 9.5-35 wt% C5-C10 alkanes, more preferably 15-25 wt%.

[0039] In this invention, the C6-C12 monocyclic alkane has the conventional meaning in the art, referring to a cycloalkane containing only one ring in its molecule, preferably selected from at least one of cyclohexane, cycloheptane, and cyclooctane.

[0040] In this invention, the monocyclic aromatic hydrocarbon has the conventional meaning in the art, referring to a hydrocarbon containing only one benzene ring, preferably selected from at least one of benzene, toluene, and xylene.

[0041] In this invention, the C5-C10 alkanes have the conventional meaning in the art, referring to alkanes without rings in their molecules, preferably selected from at least one of pentane, hexane, heptane, octane, nonane, and decane, and more preferably heptane and / or octane.

[0042] According to the present invention, the protective atmosphere, as conventionally defined in the art, refers to an atmosphere that is relatively stable and does not undergo chemical reactions during use. Preferably, the protective atmosphere is selected from at least one of nitrogen, helium, and argon.

[0043] In this invention, the equipment for the dechlorination reaction is not particularly limited, and conventional reaction apparatus in the art can be used. Preferably, the dechlorination reaction is carried out in a fixed-bed reactor.

[0044] According to a preferred embodiment of the present invention, the dechlorination reaction reactor comprises multiple fixed-bed reactors connected in parallel. Each reactor is filled with a dechlorination catalytic adsorbent, and the reactors can be switched between each other. For example, the first fixed-bed reactor is used for dechlorination initially. Once the dechlorination catalytic adsorbent in the first reactor is saturated, the reactor is switched to the next fixed-bed reactor, and so on. The dechlorination catalytic adsorbent in the saturated fixed-bed reactor can be replaced with a new one, allowing for replacement without interrupting operation and ensuring the continued progress of the overall reaction.

[0045] According to the present invention, the type and source of the dechlorination catalytic adsorbent are not particularly limited, and it can be commercially available or prepared using existing methods. According to a preferred embodiment of the present invention, the dechlorination catalytic adsorbent is selected from at least one of Al₂O₃, ZnO, CaO, and MgO.

[0046] In this invention, preferably, the dechlorination catalytic adsorbent has a porous structure, which facilitates the diffusion of petroleum resin on the surface and in the pores of the dechlorination catalytic adsorbent, thereby improving the dechlorination effect. Using the above-mentioned dechlorination catalytic adsorbent, a large amount of chlorine in the petroleum resin mixture can be converted into hydrogen chloride gas and adsorbed in a timely manner, reducing the chlorine content in the petroleum resin mixture to below 10 mg / kg.

[0047] According to the present invention, preferably, the dechlorination catalytic adsorbent contains, as oxides, Al2O3 at a mass percentage of 5-30 wt%, preferably 10-25 wt%, ZnO at a mass percentage of 0-35 wt%, preferably 0-25 wt%, CaO at a mass percentage of 0-40 wt%, preferably 15-35 wt%, and MgO at a mass percentage of 20-60 wt%, preferably 30-50 wt%.

[0048] According to the present invention, preferably, in the dechlorination reaction, the petroleum resin mixture is fed in an upflow manner.

[0049] According to the present invention, preferably, the conditions for the dechlorination reaction include: a reaction pressure of 1-3 MPa, more preferably 1.5-2.5 MPa; a reaction temperature of 150-250°C, more preferably 180-230°C; and a volume hourly space velocity of 1-6 h⁻¹. -1 Preferably 4-5 hours -1 In this invention, the dechlorination reaction is carried out under low temperature and low pressure conditions, resulting in a high chlorine removal rate, good reaction safety, and easy online switching of multiple reactors. This avoids the operational risks associated with switching reactors under high temperature and high pressure after the dechlorination catalytic adsorbent has become saturated.

[0050] According to the present invention, preferably, the conditions for the hydrogenation reaction include: a reaction temperature of 200-270°C, more preferably 200-240°C; a reaction pressure of 3-15 MPa, more preferably 5-13 MPa; and a volume hourly space velocity of 1-3 h⁻¹. -1 Preferably 1-2.5h -1 In this invention, the equipment for the hydrogenation reaction is not particularly limited, and those skilled in the art can choose suitable hydrogenation reaction equipment, preferably carrying out the hydrogenation reaction in a fixed-bed continuous reactor.

[0051] In this invention, the petroleum resin mixture enters from the bottom of the reactor and exits from the top. Nitrogen gas is introduced into the top of the reactor to stabilize the reactor pressure.

[0052] According to the present invention, preferably, in the hydrogenation reaction, the dechlorination product and hydrogen are fed in an upflow manner, i.e., the reactants enter from the bottom and exit from the top. Compared with the traditional top-in, bottom-out trickle bed reactor, this solves the problem of pressure drop during operation that forces shutdown of the hydrogenation reactor. When pressure drop occurs in the upflow reaction, it can be eliminated by backflushing the gas path, i.e., blowing gas from the top to the bottom of the reactor, without the need for shutdown, thus extending the project cycle. The upflow reactor requires less hydrogen, which can significantly reduce the hydrogen-to-oil ratio, reduce the amount of hydrogen circulating, and reduce the load on the circulating hydrogen compressor, achieving the purpose of energy saving and consumption reduction. In addition, in the hydrogenation reaction, the sulfides in the feedstock undergo hydrogenation desulfurization to generate hydrogen sulfide, which has an inhibitory effect on the activity of the catalyst. By adopting the upflow feeding mode, the hydrogen sulfide gas generated in the reactor can more easily escape from the bottom to the top of the reaction liquid zone, significantly reducing the inhibitory effect of hydrogen sulfide on the activity of the hydrogenation catalyst.

[0053] According to the present invention, preferably, the conditions for the hydrogenation reaction further include: a hydrogen-to-oil volume ratio of 50-200, for example 50, 60, 70, 80, 90, 100, 110, 120, 130, 140, 150, 160, 170, 180, or any range between the two, preferably 80-180. In the present invention, the hydrogen-to-oil volume ratio for the hydrogenation reaction refers to the volume ratio of hydrogen to petroleum resin raw material under standard conditions.

[0054] According to the present invention, preferably, the hydrogenation catalyst is an oxidized catalyst. The shape of the hydrogenation catalyst is not particularly limited, and those skilled in the art can adapt it according to the size of the catalyst bed in the hydrogenation reaction zone; preferably, it is elongated and / or spherical. Preferably, the size of the hydrogenation catalyst is slightly larger than the gaps in the Johnson mesh of the catalyst cover plate of the catalyst bed.

[0055] In this invention, preferably, the equivalent diameter of the hydrogenation catalyst is 1.3-3.5 mm, more preferably 1.5-3 mm. The equivalent diameter refers to the equivalent diameter of the radial cross-section of a strip-shaped hydrogenation catalyst or the equivalent diameter of a spherical hydrogenation catalyst.

[0056] According to the present invention, preferably, the hydrogenation catalyst comprises alumina and an active metal, the active metal comprising Group VIB metals and Group VIII metals, wherein the Group VIB metal is nickel and / or cobalt, and the Group VIII metal is molybdenum and / or tungsten.

[0057] According to the present invention, preferably, the most probable pore size of the hydrogenation catalyst is 10-20 nm, more preferably 12-18 nm. In this invention, the most probable pore size has the conventional interpretation in the art, referring to the pore size corresponding to the point of highest pore concentration. The most probable pore size of the catalyst is obtained by testing using the BET method, and the pore size corresponding to the peak value is determined by plotting the pore distribution diagram obtained from the BET method.

[0058] According to the present invention, preferably, the hydrogenation catalyst comprises secondary pores with a pore size of 100-300 nm, more preferably 150-265 nm. In the present invention, the secondary pores have the conventional meaning in the art, referring to the pores formed between particles on top of the primary pores inherent in the catalyst particles themselves. In the present invention, the hydrogenation catalyst has a secondary pore structure with a wide pore size distribution, exhibiting high tolerance to sulfur in petroleum resins. This is beneficial for the hydrogenation reaction of macromolecular petroleum resins, improving the catalyst's activity in petroleum resin hydrogenation reactions, resulting in a high heteroatom removal rate and avoiding catalyst poisoning caused by sulfur in petroleum resins.

[0059] In this invention, preferably, the pore volume of secondary pores in the hydrogenation catalyst accounts for 5-25% of the total pore volume, more preferably 10-20%. In this invention, the pore size of the secondary pores is measured by mercury porosimetry. The proportion of the pore volume of the secondary pores to the total pore volume is also measured by mercury porosimetry.

[0060] In this invention, preferably, the specific surface area of ​​the hydrogenation catalyst is 120-250 m². 2 / g, preferably 150-220m 2 / g. The pore volume of the hydrogenation catalyst is 0.4-0.9 cm³. 3 / g, preferably 0.6-0.8cm 3 / g. In this invention, the hydrogenation catalyst has a high specific surface area and pore volume, and the pore structure is controllable, which is beneficial for the loading and dispersion of active components and is suitable for the catalytic hydrogenation reaction of macromolecular petroleum resins.

[0061] In this invention, the specific surface area and pore volume were measured using the BET low-temperature nitrogen adsorption method. The experimental instrument was a Micromeritics ASAP2400 static nitrogen adsorption instrument. Experimental conditions: The sample was degassed under vacuum at 1.33 Pa and 300 °C for 4 hours, then contacted with liquid nitrogen at 77 K for isothermal adsorption and desorption. Adsorption and desorption isotherms were measured, and the specific surface area and pore volume of the catalyst were calculated using the BET formula.

[0062] According to the present invention, preferably, the mass percentage of nickel and / or cobalt, based on the total mass of the hydrogenation catalyst and in terms of oxides, is 1-15 wt%, more preferably 3-10 wt%.

[0063] According to the present invention, preferably, the mass percentage of molybdenum and / or tungsten, based on the total mass of the hydrogenation catalyst and in terms of oxides, is 8-30 wt%, more preferably 10-25 wt%.

[0064] In this invention, the content of active metal in the hydrogenation catalyst is determined by X-ray fluorescence spectroscopy.

[0065] In this invention, preferably, the hydrogenation catalyst further includes, optionally, phosphorus pentoxide, wherein the mass percentage of phosphorus pentoxide is 0-3.5 wt%.

[0066] According to the present invention, preferably, the method for preparing the hydrogenation catalyst includes:

[0067] (a) The pseudoboehmite powder is mixed with an organic additive containing multiple hydroxyl groups, shaped, and then calcined to obtain a catalyst support;

[0068] (b) Prepare a mixed solution containing an active metal source, impregnate the catalyst support in the mixed solution, dry and calcine to obtain the calcined product;

[0069] (c) The calcined product is further impregnated and dried to obtain a hydrogenation catalyst;

[0070] The active metal source includes Group VIB metals and Group VIII metals, wherein the Group VIB metals are nickel and / or cobalt, and the Group VIII metals are molybdenum and / or tungsten.

[0071] According to the present invention, preferably, the polyhydroxyl-containing organic auxiliary is selected from at least one of methylcellulose, citric acid, glucose, and ethylenediaminetetraacetic acid. In the present invention, the polyhydroxyl-containing organic auxiliary can expand the pore size of alumina, which helps to generate a secondary pore structure, and after calcination, a catalyst with a wider pore size distribution and macroporous structure is obtained.

[0072] In this invention, preferably, the organic auxiliaries containing polyhydroxyl groups are methylcellulose and / or ethylenediaminetetraacetic acid.

[0073] In this invention, the relative molecular weight of the methylcellulose is limited to a wide range. Those skilled in the art can select a suitable methylcellulose that can achieve the effects of this invention.

[0074] In this invention, preferably, the amount of the polyhydroxyl-containing organic additive added is 1-6 wt% of the mass of the boehmite powder, for example, 1 wt%, 1.5 wt%, 2 wt%, 2.5 wt%, 3 wt%, 4 wt%, 5 wt%, 6 wt%, or any range between the two, preferably 2-4 wt%.

[0075] In this invention, the mixing and molding methods are not particularly limited, as long as a uniformly mixed alumina dry strip is obtained. According to a preferred embodiment of the invention, the molding process includes: uniformly mixing boehmite powder, a polyhydroxyl-containing organic additive, optional water, optional adhesive solvent, and optional extrusion aid, and extruding to obtain alumina dry strip.

[0076] In this invention, the source of the pseudoboehmite powder is not particularly limited; it can be commercially purchased or prepared using existing methods. The pore volume of the pseudoboehmite powder is 0.8-1.3 cm³. 3 / g, specific surface area is 200-300m² 2 / g.

[0077] In this invention, the amounts of water, adhesive solvent, and extrusion aid are not particularly limited, and those skilled in the art can adjust them adaptably according to the molding conditions. The type of adhesive solvent is not particularly limited, and is a conventional adhesive solvent in the art that facilitates the molding of alumina powder, preferably selected from at least one of nitric acid, acetic acid, and citric acid. The type of extrusion aid is not particularly limited, but is preferably guar gum powder.

[0078] According to the present invention, preferably, the calcination conditions include: a calcination temperature of 500-780℃, more preferably 550-720℃; and a calcination time of 2-8h, more preferably 3-6h.

[0079] In this invention, the heating rate of calcination is not particularly limited. Preferably, the heating rate of calcination is 2-8℃ / min, more preferably 3-5℃ / min. In this invention, by controlling the calcination temperature and heating rate under the above conditions, a catalyst with a secondary pore macroporous structure can be obtained, with a more concentrated pore size distribution, which is more suitable for the hydrogenation reaction of macromolecular petroleum resins.

[0080] In this invention, preferably, the amount of catalyst support and mixed solution added is such that the mass percentage of nickel and / or cobalt in the prepared catalyst, based on the total mass of the catalyst and as oxides, is 1-15 wt%, preferably 3-10 wt%.

[0081] In this invention, preferably, the amount of catalyst support and mixed solution added is such that the mass percentage of molybdenum and / or tungsten in the prepared catalyst, based on the total mass of the catalyst and as oxides, is 8-30 wt%, preferably 10-25 wt%.

[0082] In this invention, preferably, the active metal source includes a nickel source and / or a cobalt source, as well as a molybdenum source and / or a tungsten source.

[0083] In this invention, the type of nickel source is not particularly limited, and any conventional nickel source in the art is acceptable. Preferably, the nickel source is selected from at least one of nickel acetate, nickel nitrate, nickel sulfate, and basic nickel carbonate.

[0084] In this invention, the type of molybdenum source is not particularly limited, and any conventional molybdenum source in the art is acceptable. Preferably, the molybdenum source is ammonium molybdate and / or molybdenum trioxide.

[0085] In this invention, the type of cobalt source is not particularly limited, and any conventional cobalt source in the art is acceptable. Preferably, the cobalt source is selected from at least one of cobalt sulfate, basic cobalt carbonate, and cobalt nitrate.

[0086] In this invention, the type of tungsten source is not particularly limited, and any conventional tungsten source in the art is acceptable. Preferably, the tungsten source is ammonium metatungstate and / or tungsten trioxide.

[0087] In this invention, preferably, in step (b), when the active metal source is a non-water-soluble compound, phosphoric acid and / or citric acid are added to dissolve the active metal source. The amount of phosphoric acid and / or citric acid added is not particularly limited, as long as the active metal source is fully dissolved. Those skilled in the art can adjust the amount of phosphoric acid and / or citric acid added accordingly.

[0088] In this invention, preferably, the amount of phosphoric acid added is such that the mass percentage of phosphorus pentoxide in the prepared catalyst is 0-3.5 wt% based on the total mass of the catalyst.

[0089] In this invention, the impregnation conditions are not particularly limited, and those skilled in the art can choose conventional impregnation conditions, preferably the equal volume impregnation method.

[0090] In this invention, the drying method and conditions are not particularly limited, and those skilled in the art can choose conventional drying methods and conditions. Preferably, the drying conditions include: a drying temperature of 100-200℃ and a drying time of 1-5 hours.

[0091] In this invention, preferably, the calcination conditions include: a calcination temperature of 300-480℃, more preferably 320-420℃; and a calcination time of 2-8 hours, more preferably 3-6 hours. Preferably, the calcination heating rate is 1-8℃ / min, more preferably 2-6℃ / min.

[0092] According to the present invention, preferably, the calcined product is mixed with a complexing agent solution for supplementary impregnation.

[0093] According to the present invention, preferably, the conditions for the supplementary impregnation include: an impregnation time of 1-10 hours, preferably 3-6 hours; and an impregnation temperature of 20-80°C, preferably 25-60°C.

[0094] In this invention, preferably, the complexing agent is a conventional complexing agent in the art, and is preferably selected from at least one of ethylenediaminetetraacetic acid, citric acid and glycerol.

[0095] In this invention, preferably, the mass ratio of the active metal element to the complexing agent in the calcined product is 1:0.5-10, more preferably 1:1-4. The mass concentration of the complexing agent solution is not particularly limited. According to a preferred embodiment of this invention, during the supplementary impregnation process, a fixed amount of complexing agent is weighed, and the volume of the impregnation solution is calculated based on the volume of the calcined product and its volumetric water absorption rate. The complexing agent is dissolved in deionized water to obtain an impregnation solution of a certain volume. If necessary, the solution can be heated to improve its solubility. In this invention, the supplementary impregnation product is dried to obtain a hydrogenation catalyst. The drying conditions are not particularly limited.

[0096] In this invention, supplementing the calcination product with impregnation can reduce the interaction force between the supported active metal and the support, thereby further improving the catalyst activity.

[0097] According to the present invention, preferably, the hydrogenation catalyst needs to be activated before the hydrogenation reaction. The activated hydrogenation catalyst has stronger resistance to impurities and can significantly extend its service life. The activation process is not particularly limited, but preferably, the activation process includes:

[0098] In the presence of hydrogen, the hydrogenation catalyst is sulfided using an activating oil containing an activator.

[0099] In this invention, preferably, the activation is carried out in the hydrogenation reaction zone.

[0100] According to a preferred embodiment of the present invention, activated oil and hydrogen are introduced into the hydrogenation reaction zone to form an oil circuit circulation and a gas circuit circulation, the hydrogenation catalyst is heated, and an activator is introduced for sulfidation. Preferably, the hydrogenation catalyst is heated to 140-190°C, more preferably 150-170°C, and the activator is introduced.

[0101] Preferably, the vulcanization process includes:

[0102] (1) A mixture of activated oil containing activator and hydrogenation catalyst is subjected to pressure heating to obtain a pressure-heated product;

[0103] (2) After stabilizing the product under pressure and heating, the product is heated and kept at a constant temperature to obtain an activated hydrogenation catalyst.

[0104] According to the present invention, the type of activated oil is not particularly limited and is a conventional hydrocarbon in the art. Preferably, the activated oil is selected from at least one of gasoline, kerosene, diesel, and diffusion fluid, and the source and type of the activated oil are not particularly limited. The diffusion fluid is the diffusion fluid described above, selected from at least one of C6-C12 monocyclic alkanes, monocyclic aromatics, and C5-C10 alkanes.

[0105] According to the present invention, preferably, the volume hourly space velocity of the activated oil is 1-4 h⁻¹. -1 Preferably 2-3.5h -1 .

[0106] According to the present invention, preferably, the hydrogen-oil volume ratio of the activated oil and hydrogen is 50-500, more preferably 100-300.

[0107] According to the present invention, the type and source of the activator are not particularly limited, and it can be a commonly used activator in the art, which can be commercially available or prepared by existing methods. Preferably, the activator is a sulfur-containing activator, namely dimethyl disulfide and / or FSA-55. The activator FSA-55 is a vulcanizing agent of the grade FSA-55 produced in technical cooperation with the Dalian (Fushun) Petrochemical Research Institute of Sinopec.

[0108] According to the present invention, preferably, the mass flow rate of the activator in step (1) is 0.1-0.5 wt% of the mass flow rate of the activating oil, and more preferably 0.2-0.4 wt%.

[0109] According to the present invention, preferably, the conditions for pressurization heating in step (1) include: pressurization pressure of 5-9 MPa, preferably 6-8 MPa; heating temperature of 210-250℃, preferably 220-240℃; and heating rate of 5-25℃ / h, preferably 10-20℃ / h.

[0110] According to the present invention, preferably, the stabilization conditions in step (2) include: a stabilization time of 4-12 hours, preferably 6-8 hours.

[0111] According to the present invention, preferably, the heating-holding conditions in step (2) include: heating temperature of 300-350℃, preferably 310-330℃; heating rate of 10-20℃ / h, preferably 10-15℃ / h; and holding time of 4-12h, preferably 6-8h.

[0112] According to the present invention, preferably, the hydrogen volume concentration during the activation process is not less than 60%, and more preferably 80-95%.

[0113] According to the present invention, preferably, an activator is optionally introduced in step (2).

[0114] According to the present invention, preferably, the mass flow rate ratio a% of the introduced activator to the activating oil satisfies the following condition:

[0115] H2S < 3000 mg / L, 0.3% <a≤1%;

[0116] 3000mg / L≤H2S<10000mg / L, 0.2% <a≤0.5%;

[0117] 10000mg / L≤H2S<20000mg / L, 0.1% <a≤0.2%;

[0118] 20000mg / L≤H2S, 0%≤a≤0.1%.

[0119] In this invention, an activator is introduced appropriately according to the activation status to reduce the waste of activator. After activation, the hydrogenation catalyst has good toxicity resistance and long life. Moreover, compared with precious metal catalysts, the hydrogenation catalyst of this invention is a non-precious metal catalyst with low cost and simple operation of the activation process.

[0120] Preferably, this invention further includes a step of replacing the hydrogenation reactor with a diffusion liquid after the activation process is completed. The activation oil is completely replaced, and then, with the gas path no longer circulating, fresh hydrogen is used to replace the gas path.

[0121] According to the present invention, preferably, the method further includes partially recycling the hydrogenation product back to the hydrogenation reaction zone and partially performing gas-liquid separation;

[0122] The ratio C of the amount of hydrogenation product to the amount of dechlorination product added in the recycling section satisfies:

[0123]

[0124] Where S is the sulfur content (mg / kg) in the dechlorination product, B is the bromine value (gBr / 100g) of the dechlorination product, and R is the hydrogen-to-oil volume ratio in the hydrogenation reaction.

[0125] In this invention, S represents the sulfur content (mg / kg) in the dechlorination product, which is determined using energy-dispersive X-ray fluorescence spectrometry; B represents the bromine value (gBr / 100g) of the dechlorination product, determined using the coulombic method; and R represents the hydrogen-to-oil volume ratio in the hydrogenation reaction, which refers to the volume ratio of hydrogen to petroleum resin raw material.

[0126] According to a preferred embodiment of the present invention, when the calculated value of C is less than 0.3, the value of C can be 0, that is, no recycling is required, and all hydrogenation products are separated into gas and liquid.

[0127] In this invention, the conditions and equipment for gas-liquid separation are not particularly limited, and gas-liquid separation can be carried out in conventional gas-liquid separators in the art, preferably low-pressure gas-liquid separators.

[0128] In this invention, preferably, the method further includes: performing gas-liquid separation on the hydrogenation product to obtain a liquid product, followed by flash evaporation to obtain a flash evaporated product. In this invention, the gas-liquid separation method is not particularly limited; preferably, the flash evaporation is carried out in a flash evaporator. According to a preferred embodiment of this invention, the liquid product sequentially enters a primary flash evaporator and a secondary flash evaporator for flash evaporation. Preferably, the flash evaporation pressure of the primary flash evaporator is 1-4 MPa, more preferably 1.2-2 MPa; the flash evaporation pressure of the secondary flash evaporator is 1-4 MPa, more preferably 1.2-2 MPa.

[0129] In this invention, preferably, the top stream of the flash evaporation product is recycled as a diffusion liquid.

[0130] In this invention, preferably, the bottom stream of the flash evaporation product is stripped to obtain hydrogenated petroleum resin. Preferably, the bottom product of the secondary flash evaporator enters a stripping tower for stripping to remove light component impurities from the petroleum resin. The stripping gas used in the stripping tower is selected from at least one of nitrogen, hydrogen, and methane.

[0131] In this invention, preferably, the stripping product enters the granulation unit to obtain the hydrogenated petroleum resin product.

[0132] In this invention, preferably, the method is used in conjunction with a dechlorination catalyst and a hydrogenation catalyst to process high-chlorinated petroleum resins. It achieves dechlorination of petroleum resins under low temperature and low pressure, enables online switching of multiple reactors, has high reaction safety, and reduces the Gardner color of the petroleum resin to less than 1. While achieving dechlorination and decolorization, the softening point of the petroleum resin remains largely unchanged. In this invention, the Gardner color of the petroleum resin is obtained using a colorimetric method.

[0133] The present invention will be described in detail below through examples and comparative examples. Unless otherwise specified, all reagents used in the following examples and comparative examples are commercially available.

[0134] The activator FSA-55 is a vulcanizing agent of the grade FSA-55 produced in technical cooperation with the Dalian (Fushun) Petrochemical Research Institute of Sinopec.

[0135] The boehmite powder is PB90 boehmite produced by Sinopec.

[0136] The weight-average molecular weight of the methylcellulose is 20,000 g / mol.

[0137] Example 1

[0138] Preparation of dechlorination catalytic adsorbent:

[0139] The alumina support was calcined at 600℃ for 4 hours. Zinc chloride, Ca(OH)2 and Mg(OH)2 were impregnated onto the alumina support in sequence using the equal volume impregnation method. The alumina support was then air-dried for 48 hours, dried at 130℃ for 20 hours, and calcined at 500℃ for 5 hours to obtain the dechlorination catalytic adsorbent. The composition of the dechlorination catalytic adsorbent is shown in Table 2.

[0140] Preparation of hydrogenation catalysts:

[0141] (1) Mix 100g of boehmite powder with 3g of methylcellulose and 3g of guar gum powder, add 110g of water and 3g of nitric acid and knead, extrude to form alumina carrier dry strips, place the alumina carrier dry strips in an air stream and calcine at a temperature of 720℃, a heating rate of 4℃ / min and a calcine time of 3h to obtain alumina carrier.

[0142] (2) Add a certain amount of nickel nitrate, ammonium molybdate, and ammonium metatungstate to deionized water, heat and stir to dissolve, and impregnate 50g of alumina support. Dry the impregnated catalyst at 120℃ for 2 hours, and then calcine it in an air stream for 3 hours. The calcination temperature is 350℃ and the heating rate is 2.5℃ / min to obtain the calcined product.

[0143] (3) Weigh 10g of ethylenediaminetetraacetic acid and dissolve it in water to obtain a solution. Impregnate 50g of the calcined product at 30°C for 3h to obtain a catalyst. The catalyst is named “Catalyst 1”. The active metal composition and pore structure of the catalyst 1 are shown in Table 1.

[0144] Activation status of the hydrogenation catalyst:

[0145] Activator: dimethyl disulfide; Activating oil: the diffusion liquid of Example 1.

[0146] Activation process: (1) Introduce activation oil and hydrogen (hydrogen volume concentration of 95%) into the reactor to form an oil circulation and a gas circulation. Heat to 170°C and introduce the activator. The volume hourly space velocity of the activation oil is 2 h⁻¹. -1 The mass flow rate of the activator is 0.4 wt% of the mass flow rate of the activating oil, the hydrogen-to-oil volume ratio is 100, the pressure is increased to 6 MPa, and the temperature is increased to 230℃ at 10℃ / h to obtain the heated and pressurized product.

[0147] (2) The heated and pressurized product was stabilized for 7 hours, and then the temperature was increased to 320°C at a rate of 15°C / h and held for 8 hours. The concentration of hydrogen sulfide and the mass flow rate ratio (a%) of the activator to the activating oil during this period are shown in the figure. Figure 1 ,from Figure 1 As can be seen, the value of 'a' changes with the concentration of hydrogen sulfide. Compared with Examples 2 and 3, 'a' is slightly higher in Example 1 at the same hydrogen sulfide concentration, resulting in a higher hydrogen sulfide concentration at subsequent temperatures, reaching a maximum of nearly 30,000 mg / L. However, at temperatures above 230°C, the overall hydrogen sulfide concentration is controlled between 10,000 and 30,000 mg / L, which is considered ideal.

[0148] Example 2

[0149] Preparation of dechlorination catalytic adsorbent:

[0150] The alumina support was calcined at 600℃ for 6 hours. Ca(OH)2 and Mg(OH)2 were impregnated onto the alumina support sequentially using the equal volume impregnation method. The alumina support was then air-dried for 48 hours, dried at 140℃ for 20 hours, and calcined at 500℃ for 5 hours to obtain the dechlorination catalytic adsorbent. The composition of the dechlorination catalytic adsorbent is shown in Table 2.

[0151] Preparation of hydrogenation catalysts:

[0152] (1) Mix 100g of boehmite powder with 3g of ethylenediaminetetraacetic acid (EDTA) and 3g of guar gum powder, add 105g of water and 3g of nitric acid and knead, extrude to form alumina carrier dry strips, place the alumina carrier dry strips in an air stream and calcine at a temperature of 700℃, a heating rate of 5℃ / min and a calcine time of 4h to obtain alumina carrier.

[0153] (2) Add a certain amount of ammonium molybdate and cobalt nitrate to deionized water, heat and stir to dissolve, and impregnate 50g of alumina support. Dry the impregnated catalyst at 120℃ for 2 hours, and then calcine it in an air stream for 3 hours. The calcination temperature is 400℃ and the heating rate is 2℃ / min to obtain the calcined product.

[0154] (3) Weigh 15g of ethylenediaminetetraacetic acid and dissolve it in water to obtain a solution. Impregnate 50g of the calcined product at 30°C for 3h to obtain a catalyst. The catalyst is named “catalyst 2”. The active metal composition and pore structure of the catalyst 2 are shown in Table 1.

[0155] Activation status of the hydrogenation catalyst:

[0156] Activator: FSA-55; Activating oil: Kerosene (Shanghai Petrochemical's ordinary grade 1 oil).

[0157] (1) Introduce activation oil and hydrogen (hydrogen volume concentration of 95%) into the reactor to form an oil circulation path and a gas circulation path. Heat to 150°C and introduce the activator. The volume hourly space velocity of the activation oil is 2.5 h⁻¹. -1 The mass flow rate of the activator is 0.35 wt% of the mass flow rate of the activating oil, the hydrogen-to-oil volume ratio is 110, the pressure is increased to 6 MPa, and the temperature is increased to 230℃ at 20℃ / h to obtain the heated and pressurized product.

[0158] (2) The heated and pressurized product was stabilized for 6 hours, and then the temperature was increased to 310℃ at a rate of 10℃ / h and stabilized for 8 hours. The hydrogen sulfide concentration and the mass flow rate ratio (a%) of the activator to the activating oil during this period are shown in [reference needed]. Figure 2 ,from Figure 2 It can be seen that the hydrogen sulfide concentration is basically maintained between 20,000 and 25,000 mg / L.

[0159] Example 3

[0160] Preparation of dechlorination catalytic adsorbent:

[0161] The alumina support was calcined at 600℃ for 6 hours. Zinc chloride, Ca(OH)2 and Mg(OH)2 were impregnated onto the alumina support in sequence using the equal volume impregnation method. The alumina support was then air-dried for 48 hours, dried at 130℃ for 20 hours, and calcined at 500℃ for 5 hours to obtain the dechlorination catalytic adsorbent. The composition of the dechlorination catalytic adsorbent is shown in Table 2.

[0162] Catalyst 1 from Example 1 was used as the hydrogenation catalyst.

[0163] Activation status of the hydrogenation catalyst:

[0164] Activator: dimethyl disulfide; Activated oil: diesel (Qingdao Refining & Chemical straight-run diesel).

[0165] (1) Introduce activation oil and hydrogen (hydrogen volume concentration of 95%) into the reactor to form an oil circulation path and a gas circulation path. Heat to 160°C and introduce the activator. The volume hourly space velocity of the activation oil is 3.5 h⁻¹. -1 The mass flow rate of the activator is 0.35 wt% of the mass flow rate of the activating oil, the hydrogen-to-oil volume ratio is 110, the pressure is increased to 7 MPa, and the temperature is increased to 240℃ at 20℃ / h to obtain the heated and pressurized product.

[0166] (2) The heated and pressurized product was stabilized for 6 hours, and then the temperature was increased to 330℃ at a rate of 10℃ / h and held for 8 hours. The concentration of hydrogen sulfide and the mass flow rate ratio (a%) of activator to activating oil during this period are shown in the figure. Figure 3 ,from Figure 3 It can be seen that the hydrogen sulfide concentration fluctuates within a certain range after the temperature reaches 230℃, which is related to the short duration of the high activator ratio in the early stage. However, the overall hydrogen sulfide content is between 15000-23000 mg / L.

[0167] Table 1 shows the composition and structure of catalyst 1 and catalyst 2.

[0168] Table 1

[0169]

[0170] As can be seen from Table 1, catalyst 1 and catalyst 2 have a secondary pore structure with a high proportion of secondary pores and a large pore volume. The hydrogenation catalysts loaded with active metals have high hydrogenation reaction activity and stability.

[0171] The petroleum resin dechlorination and decolorization methods in the examples and comparative examples include: using a fixed-bed reactor;

[0172] (1) In the dechlorination reaction zone, in the presence of the dechlorination catalytic adsorbent, the petroleum resin mixture is introduced by an upflow feeding method, and a protective atmosphere and diffusion liquid are introduced to carry out the dechlorination reaction and obtain the dechlorination product.

[0173] (2) The dechlorination product and hydrogen are introduced into the hydrogenation reaction zone using an upflow feed method. In the presence of the activated hydrogenation catalyst, the hydrogenation reaction takes place to obtain the hydrogenation product. A portion of the hydrogenation product undergoes gas-liquid separation to obtain a liquid product, which is then flash-evaporated to obtain the flash product. The other portion is recycled back to the hydrogenation reaction zone. The flash evaporation conditions in the examples and comparative examples include: a pressure of 1.5 MPa in the first-stage flash tank and a pressure of 1.4 MPa in the second-stage flash tank.

[0174] The properties of raw materials and products used in Examples 1, 2, and 3, as well as the catalyst and process parameters, are shown in Table 2.

[0175] Table 2

[0176]

[0177]

[0178] Note: The product nature is based on samples taken after 120 hours of continuous operation for testing.

[0179] As can be seen from Table 2, the method of the present invention for dechlorination and decolorization of petroleum resin can significantly reduce the Gardner color of petroleum resin raw materials, without significantly changing the softening point, and the dechlorination and decolorization effect is good.

[0180] Comparative Example 1

[0181] The petroleum resin was dechlorinated and decolorized according to the method in Example 1, except that no diffusion liquid was added to the petroleum resin mixture; all raw materials were petroleum resin. Other conditions were the same as in Example 1. The chlorine content of the dechlorinated product was 182 mg / kg. After hydrogenation, the Gardner color of the hydrogenated product was 2, and the softening point was 107°C.

[0182] Example 4

[0183] The petroleum resin was dechlorinated and decolorized according to the method in Example 1, except that the activator used for the hydrogenation catalyst before the hydrogenation reaction was dimethyl disulfide (DMDS). DMDS was injected at 230°C, and the value of a was maintained at 0.05 wt%. Other conditions were the same as in Example 1. After the hydrogenation reaction, the Gardner color of the hydrogenation product was 1.7, and the softening point was 107°C.

[0184] Example 5

[0185] The petroleum resin was dechlorinated and decolorized according to the method of Example 1, except that the hydrogenation catalyst used in the hydrogenation reaction was a conventional reduced hydrogenation catalyst, Ni / Al2O3 catalyst, in which the Ni content was 35wt%, and other conditions were the same as in Example 1.

[0186] After the hydrogenation reaction was carried out continuously for 2100 h, the Gardner color of the hydrogenation products of Examples 1 and 5 were <1 and 1.4, respectively, and the softening points were 105 °C and 105 °C, respectively.

[0187] As can be seen from the results in Table 1, the dechlorination and decolorization method for petroleum resins provided by this invention can effectively dechlorinate and decolorize various petroleum resins under low temperature and low pressure, significantly reducing the chlorine content in petroleum resin products and reducing Gardner color to <1.

[0188] Test case

[0189] The hydrogenation catalyst of Example 1 was subjected to a hydrogenation catalyst lifetime test experiment. The comparison sample was a commonly used reduced hydrogenation catalyst, Ni / Al2O3 catalyst, with a Ni content of 35wt%. Before using this catalyst, it needs to be reduced at 300°C and 7.0MPa in a hydrogen atmosphere.

[0190] The test used C5 petroleum resin that had not undergone dechlorination in Example 1 to directly carry out the hydrogenation reaction. Under high chlorine conditions, the deactivation of the catalyst can be accelerated, and the deactivation can be seen more clearly in the lifetime test.

[0191] The hydrogenation reaction conditions included: a reaction pressure of 7 MPa, a reaction temperature of 235 °C, a hydrogen-to-oil volume ratio of 120, and a volume hourly space velocity of 1.5 h⁻¹. -1 The results are shown in Table 3.

[0192] Table 3

[0193]

[0194] As can be seen from Table 3, the hydrogenation catalyst used in Example 1 of the present invention has a better decolorization effect and better stability than conventional hydrogenation catalysts.

[0195] The preferred embodiments of the present invention have been described in detail above; however, the present invention is not limited thereto. Within the scope of the inventive concept, various simple modifications can be made to the technical solutions of the present invention, including combinations of various technical features in any other suitable manner. These simple modifications and combinations should also be considered as the content disclosed in the present invention and are all within the protection scope of the present invention.

Claims

1. A method for dechlorinating and decolorizing petroleum resin, characterized in that, The method includes the following steps: (A) In the presence of a dechlorination catalytic adsorbent and under a protective atmosphere, a petroleum resin mixture is subjected to a dechlorination reaction to obtain a dechlorination product; (B) In the presence of a hydrogenation catalyst, the dechlorination product is mixed with hydrogen gas to carry out a hydrogenation reaction to obtain the hydrogenated product; The petroleum resin mixture includes petroleum resin raw materials and a diffusion liquid; The diffusion liquid is selected from at least one of C6-C12 monocyclic alkanes, monocyclic aromatics, and C5-C10 alkanes.

2. The method according to claim 1, wherein, In the petroleum resin mixture, the mass percentage of petroleum resin raw material is 10-40 wt%, preferably 25-35 wt%; the mass percentage of diffusion liquid is 60-90 wt%, preferably 65-75 wt%. Preferably, the diffusion liquid is selected from at least two of C6-C12 monocyclic alkanes, monocyclic aromatics, and C5-C10 alkanes; Preferably, in the diffusion liquid, the mass percentage of C6-C12 monocyclic alkanes is 65-90 wt%, preferably 75-85 wt%; the mass percentage of monocyclic aromatics is 0-0.5 wt%, preferably 0-0.3 wt%; and the mass percentage of C5-C10 alkanes is 9.5-35 wt%, preferably 15-25 wt%. Preferably, the protective atmosphere is selected from at least one of nitrogen, helium, and argon.

3. The method according to claim 1 or 2, wherein, The dechlorination catalytic adsorbent is selected from at least one of Al2O3, ZnO, CaO and MgO; Preferably, the dechlorination catalytic adsorbent contains, based on oxides, Al2O3 at a mass percentage of 5-30 wt%, preferably 10-25 wt%, ZnO at a mass percentage of 0-35 wt%, preferably 0-25 wt%, CaO at a mass percentage of 0-40 wt%, preferably 15-35 wt%, and MgO at a mass percentage of 20-60 wt%, preferably 30-50 wt%. Preferably, in the dechlorination reaction, the petroleum resin mixture is fed in an upflow manner; Preferably, the conditions for the dechlorination reaction include: a reaction pressure of 1-3 MPa, more preferably 1.5-2.5 MPa; a reaction temperature of 150-250°C, more preferably 180-230°C; and a volume hourly space velocity of 1-6 h⁻¹. -1 Preferably 4-5 hours -1 .

4. The method according to any one of claims 1-3, wherein, In the hydrogenation reaction, the dechlorination product and hydrogen are fed in an upflow manner. The conditions for the hydrogenation reaction include: a reaction temperature of 200-270℃, preferably 200-240℃; a reaction pressure of 3-15 MPa, preferably 5-13 MPa; and a volume hourly space velocity of 1-3 h⁻¹. -1 Preferably 1-2.5h -1 ; Preferably, the conditions for the hydrogenation reaction further include a hydrogen-to-oil volume ratio of 50-200, more preferably 80-180.

5. The method according to any one of claims 1-4, wherein, The hydrogenation catalyst is an oxidized catalyst; Preferably, the hydrogenation catalyst comprises alumina and an active metal, wherein the active metal comprises a Group VIB metal and a Group VIII metal, wherein the Group VIB metal is nickel and / or cobalt, and the Group VIII metal is molybdenum and / or tungsten; Preferably, the mass percentage of nickel and / or cobalt, based on the total mass of the hydrogenation catalyst and in terms of oxides, is 1-15 wt%, more preferably 3-10 wt%. Preferably, the mass percentage of molybdenum and / or tungsten, based on the total mass of the hydrogenation catalyst and in terms of oxides, is 8-30 wt%, more preferably 10-25 wt%.

6. The method according to any one of claims 1-5, wherein, The preparation method of the hydrogenation catalyst includes: (a) The pseudoboehmite powder is mixed with an organic additive containing multiple hydroxyl groups, shaped, and then calcined to obtain a catalyst support; (b) Prepare a mixed solution containing an active metal source, impregnate the catalyst support in the mixed solution, dry and calcine to obtain the calcined product; (c) The calcined product is further impregnated and dried to obtain a hydrogenation catalyst; The active metal source includes Group VIB metals and Group VIII metals, wherein the Group VIB metals are nickel and / or cobalt, and the Group VIII metals are molybdenum and / or tungsten. Preferably, the polyhydroxyl-containing organic auxiliaries are selected from at least one of methylcellulose, citric acid, glucose, and ethylenediaminetetraacetic acid; Preferably, the calcination conditions include: a calcination temperature of 500-780℃, more preferably 550-720℃; and a calcination time of 2-8h, more preferably 3-6h.

7. The method according to any one of claims 1-6, wherein, The hydrogenation catalyst needs to be activated before the hydrogenation reaction; Preferably, the activation process includes: sulfiding the hydrogenation catalyst with an activation oil containing an activator in the presence of hydrogen; Preferably, the vulcanization process includes: (1) A mixture of activated oil containing activator and hydrogenation catalyst is subjected to pressure heating to obtain a pressure-heated product; (2) After stabilizing the product under pressure and heating, the product is heated and kept at a constant temperature to obtain an activated hydrogenation catalyst.

8. The method according to claim 7, wherein, The activating oil is selected from at least one of gasoline, kerosene, diesel, and diffusion fluid; Preferably, the volume hourly space velocity (VHSV) of the activated oil is 1-4 h⁻¹. -1 Preferably 2-3.5h -1 ; Preferably, the hydrogen-to-oil volume ratio is 50-500, more preferably 100-300; Preferably, the activator is a sulfur-containing activator; Preferably, the mass flow rate of the activator in step (1) is 0.1-0.5 wt% of the mass flow rate of the activating oil, and more preferably 0.2-0.4 wt%.

9. The method according to claim 7 or 8, wherein, The conditions for pressurized heating in step (1) include: pressurization pressure of 5-9 MPa, preferably 6-8 MPa; heating temperature of 210-250℃, preferably 220-240℃; and heating rate of 5-25℃ / h, preferably 10-20℃ / h. Preferably, the stabilization conditions in step (2) include: a stabilization time of 4-12 hours, preferably 6-8 hours; Preferably, the heating-holding conditions in step (2) include: heating temperature of 300-350℃, preferably 310-330℃; heating rate of 10-20℃ / h, preferably 10-15℃ / h; and holding time of 4-12h, preferably 6-8h.

10. The method according to any one of claims 7-9, wherein, During the activation process, the hydrogen volume concentration is not less than 60%, preferably 80-95%; Preferably, an activator is optionally introduced in step (2); Preferably, the mass flow rate ratio a% of the introduced activator to the activating oil satisfies the following condition: H2S < 3000 mg / L, 0.3% <a≤1%; 3000mg / L≤H2S<10000mg / L, 0.2% <a≤0.5%; 10000mg / L≤H2S<20000mg / L, 0.1% <a≤0.2%; 20000mg / L≤H2S, 0%≤a≤0.1%.

11. The method according to any one of claims 7-10, wherein, The method also includes recycling part of the hydrogenation product back to the hydrogenation reaction zone and performing gas-liquid separation on part of it. The ratio C of the amount of hydrogenation product to the amount of dechlorination product added in the recycling section satisfies: Where S is the sulfur content (mg / kg) in the dechlorination product, B is the bromine value (gBr / 100g) of the dechlorination product, and R is the hydrogen-to-oil volume ratio in the hydrogenation reaction.

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