A combined process for the deoxygenation of a fischer-tropsch synthesis oil

Abstract

The present application relates to a combined method for deoxygenation of Fischer-Tropsch synthesis oil, which comprises that the Fischer-Tropsch synthesis oil is sequentially introduced into a hydroprotection reactor and a hydrofining reactor, and is reacted with a hydroprotection catalyst and a hydrofining catalyst to carry out olefin saturation and hydrodeoxygenation reaction, the reaction effluent of the hydrofining reactor is introduced into a make-up refining adsorption tank of an adsorption deoxygenation unit through separation to obtain a hydrogenation product oil, and is contacted with an adsorbent to carry out adsorption deoxygenation reaction, and the obtained deoxygenated hydrogenation product oil is introduced into a fractionation unit to obtain a liquid wax fraction and a wax oil fraction, thereby providing high-quality raw materials for producing high-quality liquid wax and high-value microcrystalline wax.

Description

Technical Field

[0001] The present invention relates to the technical field of Fischer-Tropsch oil processing, and in particular to a combined method for deoxygenating Fischer-Tropsch oil. Background Art

[0002] With the development of the national economy, my country's oil consumption has shown a rapid upward trend, and its dependence on foreign crude oil has increased year by year. At the same time, increasingly stringent environmental regulations have placed ever-stricter demands on the composition, impurity content, and pollutant emissions of engine fuels during processing. However, as a non-renewable energy source, oil is facing a crisis of increasing depletion, and crude oil quality is deteriorating, with a growing trend towards inferior and heavier grades. Therefore, the development of new sustainable energy technologies that meet environmental requirements and the demand for liquid fuels has become a driving force in addressing both the energy and environmental crises.

[0003] The Fischer-Tropsch synthesis reaction was discovered by German chemists Fischer and Tropsch in 1923. The reaction uses hydrogen and carbon monoxide as raw materials, producing hydrocarbons in the presence of a synthetic catalyst. The raw materials can be converted from coal, natural gas, coalbed methane, biomass, and other sources, and are widely available. The synthetic oil produced using Fischer-Tropsch synthesis technology differs significantly from conventional petroleum derivatives in terms of hydrocarbon composition and primary properties. It is primarily composed of paraffins and olefins, with a certain amount of oxygenates, with an olefin content exceeding 35%. It also has extremely low sulfur and nitrogen contents and is essentially free of cycloalkanes and aromatics. Therefore, the properties of Fischer-Tropsch synthetic oil can be exploited to remove the olefins and oxygenates through hydrogenation to produce specialty products such as light white oil and liquid wax. This offers significant advantages over producing light white oil and liquid wax from petroleum-based feedstocks.

[0004] US4500417A discloses a method for hydroconversion of Fischer-Tropsch oil, which requires presulfiding of the catalyst and injection of sulfur into the system to improve the activity and stability of the catalyst.

[0005] CN111718748A discloses a method for removing oxygenated compounds from Fischer-Tropsch oil. The oxygenated compounds in the Fischer-Tropsch oil are removed by using different extraction and distillation processes in multiple fractions. The oxygenated compound content can be reduced to below 100 ppm. The method involves multiple extraction steps.

[0006] CN112745907A relates to a method for removing oxygenates from Fischer-Tropsch oil. The method comprises liquid-liquid extraction, water washing, and solvent recovery. The extraction solvent used in the liquid-liquid extraction is at least one of an alcohol, an amide, a pyrrolidone, and a sulfone compound. The extraction solvent has high selectivity for oxygenates, good stability, and a large density difference with the Fischer-Tropsch oil.

[0007] Fischer-Tropsch oil contains a high olefin content, which releases a significant amount of heat during the hydrogenation process. Consequently, during the hydrogenation of Fischer-Tropsch oil, intense localized reactions during olefin saturation can cause excessively high bed temperatures, making them difficult to control. This can also lead to catalyst coking and carbon deposition.

[0008] In addition, Fischer-Tropsch synthetic oil basically does not contain sulfur. The use of sulfurized catalysts requires pre-sulfurization, and sulfur replenishment is required during operation to maintain the stable activity of the sulfurized catalyst. Therefore, the introduction of sulfur not only increases operating costs, but also causes a higher sulfur content in the low-fraction gas of the hydrogenation unit, increasing environmental protection costs. Summary of the Invention

[0009] The present invention aims to solve the problems of easy coking of the catalyst and high processing cost in the prior art Fischer-Tropsch synthetic oil hydrogenation process.

[0010] The present invention provides a combined method for deoxygenating Fischer-Tropsch synthetic oil, comprising:

[0011] (1) Hydrodeoxygenation unit: The Fischer-Tropsch synthesis oil is mixed with hydrogen and then enters the hydroprotection reactor and the hydrofining reactor in sequence, where it contacts the reduced hydroprotection catalyst and the reduced hydrofining catalyst to carry out olefin saturation and hydrodeoxygenation reactions. The reaction effluent of the hydrofining reactor is separated to obtain hydrogenated oil. The reaction conditions of the hydroprotection reactor are: hydrogen partial pressure of 2.0-10.0 MPa, reaction temperature of 160-290°C, hydrogen-to-oil volume ratio of 100-600, and liquid hourly volume space velocity of 4.0-28.0 h -1 The reaction conditions of the hydrotreating reactor are: hydrogen partial pressure of 2.0-10.0 MPa, reaction temperature of 290-390°C, hydrogen-to-oil volume ratio of 100-600, liquid hourly space velocity of 1.5-4.5 h -1 ;

[0012] (2) Adsorption deoxygenation unit: The hydrogenation product oil enters the supplementary refining adsorption tank, contacts with the adsorbent for adsorption deoxygenation, and obtains the deoxygenated hydrogenation product oil. The operating conditions of the supplementary refining adsorption tank are: pressure 0.05-1.0 MPa, adsorption temperature 35-150°C, liquid hourly volume space velocity 2.0-20.0 h -1 ;

[0013] (3) fractionating unit: the deoxygenated hydrogenated oil enters a fractionating system, and naphtha fraction, liquid wax fraction and wax oil fraction are obtained by fractionation, the density of the liquid wax fraction is 0.77-0.78 g / cm 3 , the oxygen content is ≤5 μg / g, the normal alkane content is 94-97 wt%, the density of the wax oil fraction is 0.80-0.82 g / cm 3 , the oxygen content is ≤5 μg / g, the normal alkane content is 92-97 wt%.

[0014] The Fischer-Tropsch synthesis oil in the present application can be a synthesis light oil, a synthesis heavy oil and a synthesis wax fraction obtained by a slurry bed Fischer-Tropsch synthesis process or a mixed fraction thereof, can be a synthesis light oil, a synthesis heavy oil and a synthesis wax fraction obtained by a fixed bed Fischer-Tropsch synthesis process or a mixed fraction thereof, or can be a mixed fraction of the Fischer-Tropsch synthesis oils obtained by the above-mentioned slurry bed process and fixed bed process.

[0015] In the present application, in order to solve the problem that the bed temperature is difficult to control due to local reaction being too violent during the process of hydrogenation saturation of olefins, the catalyst loading scheme in the hydrogenation protection reactor and the hydrogenation refining reactor and the temperature rise control scheme are preferred. Specifically, the loading volume ratio of the whole hydrogenation protection catalyst in the hydrogenation protection reactor to the whole hydrogenation refining catalyst in the hydrogenation refining reactor is 20-60:100.

[0016] The hydrogenation protection reactor is graded and loaded with three or more kinds of hydrogenation protection catalysts, and the equivalent diameter of the hydrogenation protection catalysts decreases successively and the activity increases successively along the reaction flow direction.

[0017] The total temperature rise of the hydrogenation protection reactor and the hydrogenation refining reactor is taken as a benchmark, and the temperature rise value of the hydrogenation protection reactor is controlled to be 25-70% of the total temperature rise value.

[0018] In an embodiment of the present application, the loading volume ratio of the whole hydrogenation protection catalyst in the hydrogenation protection reactor to the whole hydrogenation refining catalyst in the hydrogenation refining reactor is 20-50:100.

[0019] The total temperature rise of the hydrogenation protection reactor and the hydrogenation refining reactor is taken as a benchmark, and the temperature rise value of the hydrogenation protection reactor is controlled to be 25-60% of the total temperature rise value.

[0020] In an embodiment of the present application, the hydrogenation protection reactor is graded and loaded with three kinds of hydrogenation protection catalysts successively, namely hydrogenation protection catalyst I, hydrogenation protection catalyst II and hydrogenation protection catalyst III, and the activity of the hydrogenation protection catalysts increases successively along the reaction flow direction.

[0021] In the hydrogenation protection reactor, based on the filling volume of hydrogenation protection catalyst III, the filling volume of hydrogenation protection catalyst I is 20% to 40%, and the filling volume of hydrogenation protection catalyst II is 50% to 75%;

[0022] In one embodiment of the present invention, the equivalent diameter of the hydrogenation protection catalyst I is 9 mm to 11 mm, and its composition is 0.25% to 0.45% nickel oxide, 0.05% to 0.25% molybdenum oxide, and the balance is aluminum oxide, based on the weight of the hydrogenation protection catalyst I;

[0023] The equivalent diameter of the hydrogenation protection catalyst II is 5.6 mm to 6.5 mm. Based on the weight of the hydrogenation protection catalyst II, its composition is 0.8% to 2.8% nickel oxide, 0.15% to 0.45% molybdenum oxide, and the balance is aluminum oxide.

[0024] The equivalent diameter of the hydrogenation protection catalyst III is 2.5 mm to 3.5 mm. Based on the weight of the hydrogenation protection catalyst III, its composition is 1.8% to 3.8% nickel oxide, 0.3% to 0.75% molybdenum oxide, and the balance is aluminum oxide.

[0025] In one embodiment of the present invention, the hydroprotection reactor is graded and loaded with four hydroprotection catalysts, namely, hydroprotection catalyst I, hydroprotection catalyst II, hydroprotection catalyst III and hydroprotection catalyst IV, and the activity of the hydroprotection catalyst increases in the direction of the reactant flow;

[0026] In the hydrogenation protection reactor, based on the volume of hydrogenation protection catalyst IV, the filling volume of hydrogenation protection catalyst I is 5% to 25%, the filling volume of hydrogenation protection catalyst II is 55% to 75%, and the filling volume of hydrogenation protection catalyst III is 35% to 55%;

[0027] Preferably, the equivalent diameter of the hydrogenation protection catalyst I is 9 mm to 11 mm, and its composition, based on the weight of the hydrogenation protection catalyst I, is 0.25% to 0.45% nickel oxide, 0.05% to 0.25% molybdenum oxide, and the balance is aluminum oxide;

[0028] The equivalent diameter of the hydrogenation protection catalyst II is 5.6 mm to 6.5 mm. Based on the weight of the hydrogenation protection catalyst II, its composition is 0.8% to 2.8% nickel oxide, 0.15% to 0.45% molybdenum oxide, and the balance is aluminum oxide.

[0029] The equivalent diameter of the hydrogenation protection catalyst III is 2.5 mm to 3.5 mm. Based on the weight of the hydrogenation protection catalyst III, its composition is 1.8% to 3.8% nickel oxide, 0.3% to 0.75% molybdenum oxide, and the balance is aluminum oxide.

[0030] The equivalent diameter of the hydrogenation protection catalyst IV is 1.5mm-2.5mm, and the composition of the hydrogenation protection catalyst IV is 3.5%-7.5% of nickel oxide, 0.8%-2.8% of molybdenum oxide, and the balance of alumina, based on the weight of the hydrogenation protection catalyst IV.

[0031] In an embodiment of the present application, the hydrogenation refining catalyst is a supported catalyst, the support is alumina, the content of nickel oxide is 5.0-15.0% by weight, the content of the auxiliary element oxide is 0.8-4.5% by weight, the auxiliary element is selected from one or more of Mg, Li, Co, Fe, Mo and W, and the equivalent diameter of the hydrogenation refining catalyst is 1.3mm-1.8mm.

[0032] In an embodiment of the present application, the hydrogenation refining reactor is provided with 3-5 catalyst beds, and the hydrogenation refining catalysts filled in different beds are the same hydrogenation refining catalyst or different hydrogenation refining catalysts.

[0033] In an embodiment of the present application, the hydrogenation refining reactor is provided with 3 catalyst beds, and the catalyst filling volume of the first bed is 40-60%, and the catalyst filling volume of the second bed is 60-80%, based on the catalyst filling volume of the third bed.

[0034] The present application does not have any limitation on the preparation process of the hydrogenation protection catalyst and the hydrogenation refining catalyst.

[0035] In an embodiment of the present application, the hydrogenation refining catalyst of the present application can be obtained by the following preparation method:

[0036] The alumina shaped support is prepared from the precursor of alumina through shaping, drying, and hydrothermal stability treatment steps. The alumina can also be an industrial support. The shaping can be performed by conventional methods, such as tabletting, ball rolling, and extrusion, etc. The drying step is performed at 100-160℃ for 4-8 hours. The calcination step is performed by calcining the dried strip at 500-800℃ under air for 3-8 hours, preferably at 550-750℃ for 3-6 hours. The hydrothermal stability treatment can be steam modification, i.e. introducing water vapor in air.

[0037] The auxiliary element is selected from one or more of Mg, Li, Co, Fe, Mo and W. The auxiliary element is introduced by saturated impregnation of the support with an aqueous solution containing the auxiliary element compound, followed by drying and calcination. The drying temperature is 100-180℃, the drying time is 1-20 hours, the calcination temperature is 300-750℃, and the calcination time is 1-18 hours.

[0038] The nickel is introduced by non-saturated impregnation of an alumina support with an aqueous solution containing a nickel compound, followed by drying and optionally calcining. The drying and calcining conditions are conventional, for example, a drying temperature of 100 to 300°C, preferably 100 to 280°C, a drying time of 1 to 12 hours, preferably 2 to 8 hours; a calcination temperature of 300 to 550°C, preferably 300 to 400°C, and a calcination time of 1 to 10 hours, preferably 2 to 8 hours.

[0039] The non-saturated impregnation method refers to a method in which the ratio of the volume VL of the impregnation liquid during non-saturated impregnation to the volume Vc of the impregnation liquid during saturated impregnation satisfies the following conditions: VL / Vc=0.75-0.95, preferably VL / Vc=0.80-0.90.

[0040] In the present invention, the hydroprotection catalyst and the hydrorefining catalyst are both reduced catalysts during normal use.

[0041] In one embodiment of the present invention, the hydrogenation protected catalyst is reduced and activated in the presence of hydrogen before use to obtain the reduced hydrogenation protected catalyst. The reduction activation temperature is 340-420° C., the reduction activation time is 8-32 h, and the reduction activation operating pressure is 0.3-1.0 MPa.

[0042] In one embodiment of the present invention, the hydrorefining catalyst is reduced and activated in the presence of hydrogen before use to obtain the reduced hydrorefining catalyst. The reduction activation temperature is 340-420° C., the reduction activation time is 8-32 h, and the reduction activation operating pressure is 0.3-1.0 MPa.

[0043] In a preferred embodiment, the reduction activation of the hydrogenation protection catalyst and the hydrorefining catalyst adopts a two-step activation process, wherein the first step is a reduction activation temperature of 340-380°C and an activation time of 12-20 hours, and the second step is a reduction activation temperature of 390-420°C and an activation time of 4-12 hours;

[0044] More preferably, the first step reduction activation temperature is 350-370° C. and the activation time is 12-18 h, and the second step reduction activation temperature is 390-410° C. and the activation time is 6-12 h.

[0045] The two-step activation process can enhance the olefin hydrogenation saturation and hydrodeoxygenation activities of the hydrotreating catalyst, thereby further improving product quality.

[0046] In one embodiment of the present invention, the reaction effluent from the hydrotreating reactor undergoes gas-liquid separation to produce hydrogenated oil. The gas-liquid separators may be a combination of a hot high-pressure separator, a cold high-pressure separator, a hot low-pressure separator, and a cold low-pressure separator. The resulting hydrogenated oil then enters an adsorptive deoxygenation unit, where it contacts an adsorbent and undergoes further adsorptive deoxygenation to produce deoxygenated hydrogenated oil.

[0047] In one embodiment of the present invention, the adsorbent is selected from one or more of activated carbon, activated alumina, molecular sieve or silica gel, preferably activated carbon;

[0048] The equivalent diameter of the adsorbent is 3mm to 5mm, and the specific surface area is 600 to 1200m 2 / g, columnar in shape.

[0049] The present invention combines hydrofining and supplementary adsorption deoxygenation, and achieves deep removal of oxygen in Fischer-Tropsch oil through a simple and efficient method, and can remove the oxygen content to below 5μg / g. It provides high-quality raw materials for the production of fine chemicals and special products that require deep removal of the oxygen content of Fischer-Tropsch oil, and greatly improves the added value of Fischer-Tropsch oil.

[0050] In view of the sulfur-free characteristic of Fischer-Tropsch synthetic oil, the present invention adopts a reduced hydrogenation protection catalyst and a reduced hydrogenation refining catalyst, which can reduce the catalyst pre-sulfurization operation and the sulfur injection operation to maintain the stable activity of the sulfurized catalyst. This not only reduces the operating cost, but also improves the problem of high sulfur content in the low-gas fraction of the hydrogenation unit, thereby reducing environmental protection costs.

[0051] In addition, the present invention effectively controls the reasonable distribution of temperature rise in the hydroprotection reactor and the hydrorefining reactor by optimizing the catalyst loading ratio and grading scheme of the hydroprotection reactor and the hydrorefining reactor, avoids violent local reactions and reduces the risk of catalyst coking, and slows down the carbon deposition and deactivation of the hydrorefining catalyst, effectively improving the stable operation cycle of the reduced catalyst. BRIEF DESCRIPTION OF THE DRAWINGS

[0052] Figure 1 It is a schematic diagram of one embodiment of the combined method for deoxygenating Fischer-Tropsch synthetic oil provided by the present invention. DETAILED DESCRIPTION

[0053] The present invention will be further described below with reference to the accompanying drawings, but the present invention is not limited thereto.

[0054] Figure 1 FIG. 1 is a schematic diagram of one embodiment of the combined method for deoxygenating Fischer-Tropsch synthetic oil provided by the present invention. Figure 1As shown, the Fischer-Tropsch synthesis oil 1 is sequentially introduced into the hydrogenation protection reactor 2 and the hydrofining reactor 3 of the hydrodeoxygenation unit for reaction, the hydrofining reactor 3 reaction effluent is introduced into the hot high separator 4 for gas-liquid separation, the liquid phase from the hot high separator 4 is introduced into the hot low separator 6, the gas phase from the hot high separator 4 is introduced into the cold high separator 5 for further separation into a gas phase, liquid hydrocarbon and water 8, the gas phase separated from the cold high separator 5 is a hydrogen-rich gas, which is pressurized by a circulating compressor 15 and then mixed with new hydrogen 16 to form a circulating gas 17 which is returned to the reactor inlet and the bed layer, the liquid hydrocarbon separated from the cold high separator 5 is introduced into the cold low separator 7 together with the gas phase from the hot low separator 6 for further gas-liquid separation. The gas phase separated from the cold low separator 7 is discharged together with the gas phase 14 separated from the fractionating column. The liquid phase separated from the cold low separator 7 and the liquid phase from the hot low separator 6 are introduced into the supplemental refining adsorption tank 9 of the adsorption deoxygenation unit for adsorption deoxygenation reaction. The deoxygenated hydrodeoxygenation product from the supplemental refining adsorption tank 9 is introduced into the fractionating column 10 of the fractionating unit for separation into a naphtha fraction 11, a liquid wax fraction 12 and a wax oil fraction 13. Among them, the liquid wax fraction 12 can be used as a high-quality raw material for producing high-quality liquid wax, and the wax oil fraction 13 can be used as a high-quality raw material for producing high-value-added microcrystalline wax.

[0055] The application will be further described below in conjunction with examples, but the application should not be limited by the examples.

[0056] In the examples and comparative examples, a whole fraction of Fischer-Tropsch synthesis oil is used as the raw material, and the properties are shown in Table 1.

[0057] The hydrogenation protection catalyst I used has an equivalent diameter of 10 mm, and the composition is 0.4% of nickel oxide, 0.2% of molybdenum oxide and the balance of alumina, based on the weight of the hydrogenation protection catalyst I;

[0058] The hydrogenation protection catalyst II used has an equivalent diameter of 6 mm, and the composition is 2.0% of nickel oxide, 0.4% of molybdenum oxide and the balance of alumina, based on the weight of the hydrogenation protection catalyst II;

[0059] The hydrogenation protection catalyst III used has an equivalent diameter of 3 mm, and the composition is 3.0% of nickel oxide, 0.6% of molybdenum oxide and the balance of alumina, based on the weight of the hydrogenation protection catalyst III;

[0060] The hydrogenation protection catalyst IV used has an equivalent diameter of 2 mm, and the composition is 6.0% of nickel oxide, 2.0% of molybdenum oxide and the balance of alumina, based on the weight of the hydrogenation protection catalyst IV.

[0061] The hydrofining catalyst used has an equivalent diameter of 1.5 mm, and the composition is 8% of nickel oxide, 3% of molybdenum oxide and the balance of alumina, based on the weight of the hydrofining catalyst.

[0062] The adsorbent used is activated carbon, the equivalent diameter is 3.5mm, the specific surface area is 800m 2 / g, and the shape is columnar.

[0063] Example 1

[0064] The Fischer-Tropsch synthesis oil enters the hydrogenation protection reactor and the hydrogenation refining reactor in turn, and is reacted with the hydrogenation protection catalyst and the hydrogenation refining catalyst to carry out olefin saturation and hydrodeoxygenation reaction. The reaction effluent of the hydrogenation refining reactor is separated to obtain hydrogenation product oil which enters the make-up refining adsorption tank of the adsorption deoxygenation unit, and is reacted with the adsorbent to carry out adsorption deoxygenation reaction. The obtained deoxygenated hydrogenation product oil enters the fractionation unit to obtain liquid wax fraction and wax oil fraction, which provides high-quality raw materials for producing high-quality liquid wax and high-value microcrystalline wax.

[0065] The volume ratio of the hydrogenation protection catalyst to the hydrogenation refining catalyst is 50:100.

[0066] In the hydrogenation protection reactor, the loading volume of the hydrogenation protection catalyst I is 16%, the loading volume of the hydrogenation protection catalyst II is 65%, and the loading volume of the hydrogenation protection catalyst III is 45%, based on the loading volume of the hydrogenation protection catalyst IV.

[0067] The hydrogenation refining reactor is provided with three catalyst beds. The loading volume of the catalyst of the first bed is 50%, and the loading volume of the catalyst of the second bed is 70%, based on the loading volume of the catalyst of the third bed. The hydrogenation refining catalysts loaded in the three beds are the same catalyst.

[0068] The hydrogenation protection catalyst and the hydrogenation refining catalyst are reduced and activated in the presence of hydrogen before use to obtain reduced hydrogenation protection catalyst and reduced hydrogenation refining catalyst, respectively. The reduction and activation is carried out in two steps. The first step is reduction and activation at a temperature of 360℃ for 18h. The second step is reduction and activation at a temperature of 400℃ for 6h. The specific operation conditions, reaction performance and product properties are shown in Tables 2 and 3.

[0069] As can be seen from Table 3, the olefin saturation rate of the deoxygenated hydrogenation product oil is 99.1%, and the deoxygenation rate is 99.95%, based on the raw oil. In terms of reaction performance, due to the reasonable grading of the catalysts in the hydrogenation protection reactor and the hydrogenation refining reactor, the temperature rise of the hydrogenation protection reactor and the hydrogenation refining reactor is 26℃ and 32℃, respectively, which ensures that the reaction heat is reasonably distributed in the two reactors. After 8300 hours of operation, the device is stopped and the catalyst is unloaded. After unloading, the hydrogenation refining catalyst is detected, and the carbon content is measured to be 3.6wt%. It shows that the hydrogenation refining catalyst has less carbon deposition and no coking phenomenon, which ensures the long-period stable operation of the device.

[0070] As can be seen from Table 3, the oxygen content of the produced liquid wax fraction is 1.03 μg / g, which achieves the purpose of deep deoxygenation and can be used as a high-quality raw material for producing high-quality liquid wax; the oxygen content of the wax oil fraction is 2.35 μg / g, which also achieves the purpose of deep deoxygenation and can be used as a high-quality raw material for producing high-value-added microcrystalline wax.

[0071] Example 2

[0072] The Fischer-Tropsch synthetic oil enters the hydroprotection reactor and the hydrotreating reactor in sequence, contacts and reacts with the hydroprotection catalyst and the hydrotreating catalyst to carry out olefin saturation and hydrodeoxygenation reactions. The reaction effluent of the hydrotreating reactor is separated to obtain the hydrogenated oil which enters the supplementary refining adsorption tank of the adsorption deoxygenation unit, contacts with the adsorbent to carry out adsorption deoxygenation reaction. The obtained deoxygenated hydrogenated oil enters the fractionation unit to obtain liquid wax fraction and wax oil fraction, providing high-quality raw materials for the production of high-quality liquid wax and high-value-added microcrystalline wax.

[0073] The volume ratio of the hydroprotection catalyst to the hydrorefining catalyst is 37.5:100.

[0074] In the hydrogenation protection reactor, based on the loading volume of hydrogenation protection catalyst IV, the loading volume of hydrogenation protection catalyst I in the hydrogenation protection reactor is 5%, the loading volume of hydrogenation protection catalyst II is 55%, and the loading volume of hydrogenation protection catalyst III is 35%.

[0075] The hydrotreating reactor is equipped with three catalyst beds. Based on the catalyst loading volume of the third bed, the catalyst loading volume of the first bed is 40%, and the catalyst loading volume of the second bed is 60%. The hydrotreating catalyst loaded in the three beds is the same catalyst.

[0076] Before use, the hydroprotected catalyst and hydrorefining catalyst were reduced and activated in the presence of hydrogen to obtain the reduced hydroprotected catalyst and reduced hydrorefining catalyst, respectively. The reduction and activation process was performed in a stepwise manner: the first step was at a temperature of 345°C for 12 hours; the second step was at a temperature of 415°C for 12 hours. Specific operating conditions, reaction performance, and product properties are shown in Tables 2 and 3.

[0077] As can be seen from Table 3, the olefin saturation rate of the deoxygenated hydrogenated product oil is 98.5% and the deoxygenation rate is 99.91% based on the raw oil by using the method combining hydrodeoxygenation and adsorption deoxygenation. The oxygen content of the liquid wax fraction produced is 4.87 μg / g, achieving the purpose of deep deoxygenation and the liquid wax fraction can be used as a high-quality raw material for producing high-quality liquid wax. The oxygen content of the wax oil fraction is 4.98 μg / g, also achieving the purpose of deep deoxygenation and the wax oil fraction can be used as a high-quality raw material for producing high-value microcrystalline wax. In terms of reaction performance, due to the reasonable grading of the catalysts in the hydrogenation protection reactor and the hydrogenation refining reactor, the temperature rise of the hydrogenation protection reactor and the hydrogenation refining reactor is 22°C and 36°C respectively, which ensures that the reaction heat is reasonably distributed in the two reactors. The device is stopped for unloading after 8600 hours of operation, and the hydrogenation refining catalyst is detected after unloading. The carbon content is 4.8 wt%, indicating that the hydrogenation refining catalyst has less carbon deposition and no coking phenomenon, ensuring the long-term stable operation of the device.

[0078] Example 3

[0079] The Fischer-Tropsch synthesis oil enters the hydrogenation protection reactor and the hydrogenation refining reactor in turn, and is reacted with the hydrogenation protection catalyst and the hydrogenation refining catalyst to perform olefin saturation and hydrodeoxygenation reaction. The reaction effluent of the hydrogenation refining reactor is separated to obtain hydrogenated product oil which enters the make-up refining adsorption tank of the adsorption deoxygenation unit and is contacted with the adsorbent to perform adsorption deoxygenation reaction. The deoxygenated hydrogenated product oil enters the fractionation unit to obtain liquid wax fraction and wax oil fraction, providing high-quality raw materials for producing high-quality liquid wax and high-value microcrystalline wax.

[0080] The volume ratio of the hydrogenation protection catalyst to the hydrogenation refining catalyst is 25:100.

[0081] In the hydrogenation protection reactor, the loading volume of the hydrogenation protection catalyst I is 25%, the loading volume of the hydrogenation protection catalyst II is 75%, and the loading volume of the hydrogenation protection catalyst III is 55% based on the loading volume of the hydrogenation protection catalyst IV.

[0082] The hydrogenation refining reactor is provided with three catalyst beds. The loading volume of the catalyst in the first bed is 60% and the loading volume of the catalyst in the second bed is 80% based on the loading volume of the catalyst in the third bed. The hydrogenation refining catalysts loaded in the three beds are the same catalyst.

[0083] The hydroprotected catalyst and hydrorefining catalyst were reduced and activated in the presence of hydrogen to obtain the reduced hydroprotected catalyst and reduced hydrorefining catalyst, respectively. The reduction and activation was performed in a stepwise manner: the first step was at 380°C for 20 hours; the second step was at 390°C for 4 hours. Specific operating conditions, reaction performance, and product properties are shown in Tables 2 and 3.

[0084] As shown in Table 3, the combined hydrodeoxygenation and adsorption deoxygenation method achieves an olefin saturation rate of 99.6% and a deoxygenation rate of 99.98% for the oil produced by the deoxygenation hydrogenation process, based on the feedstock. The oxygen content of the liquid wax fraction produced is 3.24 μg / g, demonstrating deep deoxygenation and suitable for use as a high-quality feedstock for producing high-quality liquid wax. The oxygen content of the wax oil fraction is 3.76 μg / g, also demonstrating deep deoxygenation and suitable for use as a high-quality feedstock for producing high-value-added microcrystalline wax. Regarding reaction performance, due to the optimal catalyst gradation in the guard and refining reactors, the temperature rises in the guard and refining reactors are 32°C and 34°C, respectively, ensuring optimal distribution of reaction heat between the two reactors.

[0085] Example 4

[0086] The Fischer-Tropsch synthetic oil enters the hydroprotection reactor and the hydrotreating reactor in sequence, contacts and reacts with the hydroprotection catalyst and the hydrotreating catalyst to carry out olefin saturation and hydrodeoxygenation reactions. The reaction effluent of the hydrotreating reactor is separated to obtain the hydrogenated oil which enters the supplementary refining adsorption tank of the adsorption deoxygenation unit, contacts with the adsorbent to carry out adsorption deoxygenation reaction. The obtained deoxygenated hydrogenated oil enters the fractionation unit to obtain liquid wax fraction and wax oil fraction, providing high-quality raw materials for the production of high-quality liquid wax and high-value-added microcrystalline wax.

[0087] The volume ratio of the hydroprotection catalyst to the hydrorefining catalyst is 45:100.

[0088] The hydroprotection reactor is loaded with three hydroprotection catalysts in a graded manner: hydroprotection catalyst I, hydroprotection catalyst II, and hydroprotection catalyst III. Based on the loading volume of hydroprotection catalyst III, the loading volume of hydroprotection catalyst I in the hydroprotection reactor is 30%, and the loading volume of hydroprotection catalyst II is 65%.

[0089] The hydrotreating reactor is equipped with three catalyst beds. Based on the catalyst loading volume of the third bed, the catalyst loading volume of the first bed is 60%, and the catalyst loading volume of the second bed is 80%. The hydrotreating catalyst loaded in the three beds is the same catalyst.

[0090] Before use, the hydroprotected catalyst and hydrorefining catalyst were reduced and activated in the presence of hydrogen to obtain the reduced hydroprotected catalyst and reduced hydrorefining catalyst, respectively. The reduction and activation process was performed in a step-by-step manner: the first step was at 380°C for 20 hours; the second step was at 390°C for 4 hours. Specific operating conditions, reaction performance, and product properties are shown in Tables 2 and 3.

[0091] As shown in Table 3, the combined hydrodeoxygenation and adsorption deoxygenation method achieves an olefin saturation rate of 99.4% and a deoxygenation rate of 99.96% for the deoxygenated hydrogenated oil, based on the feedstock. The oxygen content of the produced liquid wax fraction is 3.43 μg / g, demonstrating deep deoxygenation and suitable for use as a high-quality feedstock for producing high-quality liquid wax. The oxygen content of the wax oil fraction is 3.88 μg / g, also demonstrating deep deoxygenation and suitable for use as a high-quality feedstock for producing high-value-added microcrystalline wax. In terms of reaction performance, due to the optimal catalyst gradation in the hydroguard reactor and hydrofinishing reactor, the temperature rises in the hydroguard reactor and hydrofinishing reactor were 31°C and 35°C, respectively, ensuring optimal distribution of reaction heat between the two reactors.

[0092] Example 5

[0093] This embodiment adopts the same process distillation range as that of Example 1, the same catalyst grading scheme for the hydroprotection reactor and the same catalyst grading scheme for the hydrofining reactor, and the same adsorbent.

[0094] Example 5 differs from Example 1 in that the reduction activation process for the hydroprotected catalyst and the hydrorefining catalyst utilizes a one-step reduction process. During the reduction activation, the temperature is directly raised from 30°C to 360°C, eliminating the 6-hour activation step at 400°C. The total reduction activation time at 360°C remains the same as in Example 1, at 24 hours. Specific operating conditions, reaction performance, and product properties are shown in Tables 2 and 3.

[0095] As can be seen from Table 3, Example 5 employs a one-step reduction process. Based on the feedstock oil, the olefin saturation rate of the hydrotreated oil is 97.6%, and the deoxygenation rate of the hydrotreated oil is 91.44%. The hydrorefining effect of the Fischer-Tropsch synthesis oil is inferior to that of Example 1. To ensure the ultimate deoxygenation effect, the operating temperature of the adsorption deoxygenation is appropriately increased in Example 5. As a result, after adsorption deoxygenation and fractionation, the oxygen content of the liquid wax fraction produced is 4.11 μg / g, and the oxygen content of the wax oil fraction is 4.88 μg / g. Both are higher than the oxygen contents of the liquid wax fraction and the wax oil fraction obtained in Example 1, and the product still provides a high-quality raw material for producing high-quality liquid wax and high-value-added microcrystalline wax.

[0096] Comparative Example 1

[0097] Comparative Example 1 differs from Example 1 in that only hydrodeoxygenation was used in Comparative Example 1, without adsorption deoxygenation. Specific operating conditions, reaction performance, and product properties are shown in Tables 2 and 3.

[0098] As can be seen from Table 3, Comparative Example 1 used only hydrodeoxygenation to remove oxygenated compounds. Based on the feedstock oil, the deoxygenation rate of the hydrodeoxygenated oil was 97.46%, which was inferior to that of Example 1. The oxygen content of the produced liquid wax fraction was 78 μg / g, failing to achieve deep deoxygenation; the oxygen content of the wax oil fraction was 113 μg / g, also failing to achieve deep deoxygenation. Because the produced liquid wax and wax oil fractions were not deeply deoxygenated, they are not suitable as feedstock for the production of high-quality liquid wax and high-value-added microcrystalline wax.

[0099] Comparative Example 2

[0100] This comparative example uses the same process distillation range as Example 1, except that the volume ratio of the hydrogenation protection catalyst to the hydrorefining catalyst in Comparative Example 3 is 10:100. Specific operating conditions, reaction performance, and product properties are shown in Tables 2 and 3.

[0101] As shown in Table 3, the temperature rises in the hydroguard reactor and hydrofining reactor in Comparative Example 2 were 8°C and 50°C, respectively. The reaction exotherm was too concentrated in the hydrofining reactor, preventing proper distribution of the reaction heat between the two reactors. After 5160 hours of operation, the unit was shut down for catalyst unloading. The carbon content on the unloaded hydrofining catalyst was measured to be 16.6% by weight. This indicates that the concentrated reaction heat resulted in significant carbon deposition on the hydrofining catalyst, with varying degrees of coking occurring in both the first and second beds of the hydrofining reactor, significantly shortening the unit's operating time.

[0102] Table 1 Properties of crude oil

[0103]

[0104]

[0105] Table 2 Operating conditions

[0106]

[0107]

[0108] Table 3 Reaction performance evaluation results and product properties

[0109]

[0110]

[0111]

Claims

1. A combined method for deoxygenating a Fischer-Tropsch oil, comprising: (1) Hydrodeoxygenation unit: The Fischer-Tropsch synthesis oil is mixed with hydrogen and then enters the hydroprotection reactor and the hydrofining reactor in sequence, where it contacts the reduced hydroprotection catalyst and the reduced hydrofining catalyst to carry out olefin saturation and hydrodeoxygenation reactions. The reaction effluent of the hydrofining reactor is separated to obtain hydrogenated oil. The reaction conditions of the hydroprotection reactor are: hydrogen partial pressure of 2.0-10.0 MPa, reaction temperature of 160-290°C, hydrogen-to-oil volume ratio of 100-600, and liquid hourly volume space velocity of 4.0-28.0 h -1 The reaction conditions of the hydrotreating reactor are: hydrogen partial pressure of 2.0-10.0 MPa, reaction temperature of 290-390°C, hydrogen-to-oil volume ratio of 100-600, liquid hourly space velocity of 1.5-4.5 h -1 Taking the total temperature rise of the hydroprotection reactor and the hydrofining reactor as the benchmark, the temperature rise of the hydroprotection reactor is controlled to be 25-70% of the total temperature rise; (2) Adsorption deoxygenation unit: The hydrogenation product oil enters the supplementary refining adsorption tank, contacts with the adsorbent for adsorption deoxygenation, and obtains the deoxygenated hydrogenation product oil. The operating conditions of the supplementary refining adsorption tank are: pressure 0.05-1.0 MPa, adsorption temperature 35-150°C, liquid hourly volume space velocity 2.0-20.0 h -1 ; (3) Fractionation unit: The deoxygenated hydrogenated oil enters the fractionation system and is fractionated to obtain naphtha fraction, liquid wax fraction and wax oil fraction. The density of the liquid wax fraction is 0.77-0.78 g / cm 3 , oxygen content ≯5μg / g, normal alkane content 94-97% by weight, and density of the wax oil fraction 0.80-0.82g / cm 3 , oxygen content ≯5μg / g, normal alkane content 92~97% by weight.

2. The method according to claim 1, characterized in that The loading volume ratio of the integral hydroprotectant catalyst in the hydroprotection reactor to the integral hydrorefining catalyst in the hydrorefining reactor is 20 to 60:100; The hydrogenation protection reactor is graded and loaded with three or more hydrogenation protection catalysts. The equivalent diameters of the hydrogenation protection catalysts decrease and the activities increase in sequence along the direction of the reactant flow.

3. The method according to claim 2, characterized in that The loading volume ratio of the integral hydroprotectant catalyst in the hydroprotection reactor to the integral hydrorefining catalyst in the hydrorefining reactor is 20 to 50:100; Taking the total temperature rise of the hydrogenation protection reactor and the hydrorefining reactor as a benchmark, the temperature rise of the hydrogenation protection reactor is controlled to be 25-60% of the total temperature rise.

4. The method according to claim 1 or 2, characterized in that The hydrogenation protection reactor is graded and loaded with three hydrogenation protection catalysts, namely hydrogenation protection catalyst I, hydrogenation protection catalyst II and hydrogenation protection catalyst III, and the activity of the hydrogenation protection catalyst increases in the direction of the reactant flow; In the hydrogenation protection reactor, based on the filling volume of hydrogenation protection catalyst III, the filling volume of hydrogenation protection catalyst I is 20% to 40%, and the filling volume of hydrogenation protection catalyst II is 50% to 75%.

5. The method according to claim 4, characterized in that The equivalent diameter of the hydrogenation protection catalyst I is 9 mm to 11 mm. Based on the weight of the hydrogenation protection catalyst I, its composition is 0.25% to 0.45% nickel oxide, 0.05% to 0.25% molybdenum oxide, and the balance is aluminum oxide. The equivalent diameter of the hydrogenation protection catalyst II is 5.6 mm to 6.5 mm. Based on the weight of the hydrogenation protection catalyst II, its composition is 0.8% to 2.8% nickel oxide, 0.15% to 0.45% molybdenum oxide, and the balance is aluminum oxide. The equivalent diameter of the hydrogenation protection catalyst III is 2.5 mm to 3.5 mm. Based on the weight of the hydrogenation protection catalyst III, its composition is 1.8% to 3.8% nickel oxide, 0.3% to 0.75% molybdenum oxide, and the balance is aluminum oxide.

6. The method according to claim 1 or 2, characterized in that The hydrogenation protection reactor is graded and loaded with four hydrogenation protection catalysts, namely hydrogenation protection catalyst I, hydrogenation protection catalyst II, hydrogenation protection catalyst III and hydrogenation protection catalyst IV, and the activity of the hydrogenation protection catalyst increases in the direction of the reactant flow; In the hydrogenation protection reactor, based on the filling volume of hydrogenation protection catalyst IV, the filling volume of hydrogenation protection catalyst I is 5% to 25%, the filling volume of hydrogenation protection catalyst II is 55% to 75%, and the filling volume of hydrogenation protection catalyst III is 35% to 55%.

7. The method according to claim 6, characterized in that The equivalent diameter of the hydrogenation protection catalyst I is 9 mm to 11 mm. Based on the weight of the hydrogenation protection catalyst I, its composition is 0.25% to 0.45% nickel oxide, 0.05% to 0.25% molybdenum oxide, and the balance is aluminum oxide. The equivalent diameter of the hydrogenation protection catalyst II is 5.6 mm to 6.5 mm. Based on the weight of the hydrogenation protection catalyst II, its composition is 0.8% to 2.8% nickel oxide, 0.15% to 0.45% molybdenum oxide, and the balance is aluminum oxide. The equivalent diameter of the hydrogenation protection catalyst III is 2.5 mm to 3.5 mm. Based on the weight of the hydrogenation protection catalyst III, its composition is 1.8% to 3.8% nickel oxide, 0.3% to 0.75% molybdenum oxide, and the balance is aluminum oxide. The equivalent diameter of the hydrogenation protection catalyst IV is 1.5 mm to 2.5 mm. Based on the weight of the hydrogenation protection catalyst IV, its composition is 3.5% to 7.5% nickel oxide, 0.8% to 2.8% molybdenum oxide, and the balance is aluminum oxide.

8. The method according to claim 1, characterized in that The hydrorefining catalyst is a supported catalyst, the carrier is alumina, and based on the hydrorefining catalyst, the content of nickel oxide is 5.0 to 15.0% by weight, the content of the oxide of the auxiliary element is 0.8 to 4.5% by weight, and the auxiliary element is selected from one or more of Mg, Li, Co, Fe, Mo and W. The equivalent diameter of the hydrorefining catalyst is 1.3 mm to 1.8 mm.

9. The method according to claim 1, characterized in that The hydrorefining reactor is provided with 3 to 5 catalyst beds, and the hydrorefining catalysts loaded in different beds are the same hydrorefining catalyst or different hydrorefining catalysts.

10. The method according to claim 1 or 9, characterized in that The hydrofining reactor is provided with three catalyst beds. Taking the catalyst filling volume of the third bed as a basis, the catalyst filling volume of the first bed is 40-60%, and the catalyst filling volume of the second bed is 60-80%.

11. The method according to claim 1, wherein Before use, the hydrogenation protection catalyst is reduced and activated in the presence of hydrogen to obtain the reduced hydrogenation protection catalyst. The reduction activation temperature is 340-420° C., the reduction activation time is 8-32 hours, and the reduction activation operating pressure is 0.3-1.0 MPa. Before use, the hydrorefining catalyst is reduced and activated in the presence of hydrogen to obtain the reduced hydrorefining catalyst. The reduction activation temperature is 340-420° C., the reduction activation time is 8-32 hours, and the reduction activation operating pressure is 0.3-1.0 MPa.

12. The method according to claim 11, characterized in that The reduction activation of the hydrogenation protection catalyst and the hydrorefining catalyst adopts a two-step activation process. The first step is a reduction activation temperature of 340-380°C and an activation time of 12-20 hours. The second step is a reduction activation temperature of 390-420°C and an activation time of 4-12 hours.

13. The method according to claim 12, characterized in that The first step of reduction activation temperature is 350-370° C. and the activation time is 12-18 h. The second step of reduction activation temperature is 390-410° C. and the activation time is 6-12 h.

14. The method according to claim 1, wherein The adsorbent is selected from one or more of activated carbon, activated alumina, molecular sieve or silica gel; The equivalent diameter of the adsorbent is 3mm to 5mm, and the specific surface area is 600 to 1200m 2 / g, columnar in shape.

15. The method according to claim 1, wherein The adsorbent is activated carbon.

Images