Process for producing fuels and / or chemicals from renewable resources

Abstract

A process for hydrotreating a renewable feedstock includes passing the renewable feedstock to a first catalyst zone comprising a single metal Group VIB catalyst under hydrodeoxygenation conditions to produce a reaction mixture comprising a hydrotreated oil having an olefin content at least 250 times greater than the thermodynamic equilibrium olefin content. The reaction mixture is then passed to a second catalyst zone comprising a bimetallic catalyst having a Group VIII metal and a Group VIB metal under hydrogenation conditions to produce an effluent comprising a hydrotreated oil having an olefin content close to the thermodynamic equilibrium olefin content.

Description

Technical Field

[0001] This invention relates to the field of producing fuels and / or chemicals from renewable resources, and more particularly to methods for increasing the yield of fuels and / or chemicals from renewable resources. Background Technology

[0002] The increased energy demand resulting from global economic growth and development has led to an increase in the concentration of greenhouse gases in the atmosphere. This is considered one of the most important challenges facing humanity in the 21st century. To mitigate the impact of greenhouse gases, efforts have been made to reduce the global carbon footprint. The Earth system's capacity to absorb greenhouse gas emissions has been exhausted. Therefore, the goal is to achieve net-zero emissions by 2050. To achieve these emission reduction targets, the world is transitioning from a purely traditional carbon-based fossil fuel energy source. Achieving this energy transition in a timely manner requires a multi-pronged approach, including, for example, energy conservation, improved energy efficiency, electrification, and efforts to utilize renewable resources to produce fuels and fuel components and / or chemical feedstocks.

[0003] For example, vegetable oils, oils derived from algae, and animal fats are considered new sources for low-carbon fuel production. Similarly, deconstructed materials are seen as potential sources of low-carbon renewable fuel materials, such as pyrolyzed recyclable materials or wood.

[0004] Renewable feedstocks can include materials with very high molecular weights and high viscosity, such as triglycerides, meaning that using them directly or as a mixture as fuel base is problematic for modern engines. On the other hand, the hydrocarbon chains that make up, for example, triglycerides are essentially linear, and their length (in terms of carbon number) is compatible with hydrocarbons used / used as fuel. Therefore, converting feedstocks containing triglycerides to obtain high-quality fuel components is attractive. Similarly, renewable feedstocks can contain unsaturated compounds and / or oxygenated compounds as unsaturated compounds.

[0005] Therefore, renewable feedstocks are hydrogenated to remove oxygen, sulfur, nitrogen, metals, and olefins.

[0006] Bimetallic catalysts, typically containing Group VIII and Group VIB metals, are used for hydrogenation processes. Of particular note are NiMo, NiW, and CoMo catalysts. A challenge with bimetallic catalysts is that the yield of alkanes is affected by the generation of carbon oxides (CO and CO2). Furthermore, these carbon oxide byproducts must be treated to avoid potentially harmful environmental impacts.

[0007] Another option is to use a single-metal Group VIB metal, which has been shown to reduce the production levels of carbon oxides. This reduction in carbon oxide production leads to a decrease in the carbon loss of desired alkanes, such as C. 16 and C 18Alkanes, not C 15 and C 17 Alkanes.

[0008] An example of a single-metal Group VIB metal catalyst is described by Egenberg et al. (US8,912,375B2, December 16, 2014). In Egenberg et al., fossil-based feedstocks were co-processed with 1% to 35% renewable feedstock using unpromoted Mo catalysts containing 0.1% to 20% Mo. Egenberg et al. found that the unpromoted catalyst impregnated only with Mo (instead of Co or Ni) was effective in converting vegetable oils, animal fats, fatty acids, and fatty acid methyl esters into n-alkanes. As explained therein, the unpromoted catalyst exhibited slightly lower activity than conventional NiMo or CoMo catalysts. This reduced the likelihood of coke formation and other deposits that impair catalyst activity and lead to increased pressure drop. By using a support with a bimodal pore distribution, the catalyst employed in this method was more resistant to pore blockage and minimized increases in pressure drop and deactivation rate.

[0009] Bhan (US11,434,431B2, September 6, 2022) provides an improved catalyst having 5 wt% to 25 wt% Mo supported on a silica-alumina support having an average pore size of 40 Å to 500 Å and a pore size of 200 μm. 2 / g to 500m 2 / g nitrogen surface area. The bio-derived feedstock is contacted with the catalyst to produce a conversion product having an oxygen content reduced to below the initial oxygen content of the feedstock.

[0010] The challenge of using Mo-only catalysts to hydrogenate renewable feedstocks is that Mo-only catalysts have low hydrogenation strength.

[0011] Efforts have been made to find catalyst systems that address the challenges of monometallic and bimetallic catalyst systems.

[0012] Abhari et al. (US8,026,401B2, September 27, 2011) described a hydrodeoxygenation process in two selective steps. The first step uses a low-activity catalyst, such as Haldor Topsoe TK-709, to crack triglycerides in the feed into fatty alcohols and diesters. The second step requires a higher-activity catalyst, such as NiMo, to hydrodeoxygenate the fatty alcohols and diesters into alkanes. The introduction of the NiMo catalyst in the hydrotreatment of the organic oxygen-containing compounds (i.e., fatty alcohols and diesters) leads to the formation of carbon oxides.

[0013] Gomes et al. (US8,507,738B2, August 13, 2013) relates to the hydrotreating of feedstocks containing between 1% and 75% by weight of biomass oil diluted in a refinery petroleum hydrocarbon stream. The feedstock is contacted with a first monometallic catalytic bed and at least one second bimetallic catalytic bed, the first monometallic catalytic bed containing only one Group VIB metal oxide, selectively used to hydroconvert the biomass oil to C. 12 -C 18 This second bimetallic catalytic bed is used for the hydrogenation of unsaturated hydrocarbons and the removal of heteroatom contaminants from refinery petroleum hydrocarbon streams. An example in Gomes et al. shows an invention using a first bed with 50 vol% of a monometallic Mo catalyst and a second bed with a bimetallic NiMo catalyst.

[0014] There is still a need to improve catalyst activity and the yield of desired alkanes for fuel and / or chemical production. Summary of the Invention

[0015] According to one aspect of the present invention, a method for hydrogenating a renewable feedstock is provided, the method comprising the steps of: providing a renewable feedstock; passing the renewable feedstock under hydrodeoxygenation conditions to a first catalyst zone comprising a single-metal Group VIB catalyst to produce a first reaction mixture comprising a hydrotreated oil having an olefin content at least 250 times greater than the thermodynamic equilibrium olefin content; and passing the first reaction mixture under hydrogenation conditions to a second catalyst zone comprising a bimetallic catalyst having a Group VIII metal and a Group VIB metal to produce an effluent comprising a hydrotreated oil having an olefin content close to the thermodynamic equilibrium olefin content. Attached Figure Description

[0016] The method of the present invention will be better understood by referring to the following preferred embodiments and the accompanying drawings referenced therein, wherein:

[0017] Figures 1 to 5 This is a flowchart illustrating an embodiment of the method of the present invention. Detailed Implementation

[0018] The present invention provides a method for hydrogenating renewable feedstocks using a catalyst system having a first catalyst zone comprising a monometallic catalyst having a Group VIB metal and a second catalyst zone comprising a bimetallic catalyst having a Group VIII metal and a Group VIB metal.

[0019] Contrary to conventional wisdom, which holds that hydrotreating / deoxygenation reactions require catalysts with sufficiently strong hydrogenation capabilities, rendering single-metal catalysts unsuitable, this invention utilizes the negative characteristics of conventionally understood single-metal catalysts for olefin production. This invention provides the complete conversion of organic oxygen in a regenerable feedstock within a first catalyst zone, thereby producing olefins that are generally considered harmful. The advantage of complete organic oxygen conversion in the first catalyst zone is that C... 16 and C 18 Hydrocarbons are preserved with almost no carbon production. 15 and C 17 Hydrocarbons. The reaction mixture from the first catalyst zone contains a hydrotreated oil with an olefin content exceeding the thermodynamic equilibrium olefin content by at least 250 times, preferably at least 500 times, and more preferably at least 1000 times. Preferably, the hydrotreated oil has a reference olefin content of at least 125 ppm-w, preferably 250 ppm-w, and more preferably 500 ppm-w.

[0020] Hydrogenation catalysts catalyze hydrogenation and dehydrogenation reactions. Under typical operating pressures and temperatures, thermodynamics indicates that the equilibrium reaction favors the alkane component almost entirely, with only trace amounts (e.g., a few ppm-w or less) of the olefin component in the equilibrium mixture. The strength of the hydrogenation / dehydrogenation function of the hydrogenation catalyst and the available space velocity determine whether the reaction rate is sufficient to reach thermodynamic equilibrium between the alkane and olefin components.

[0021] Olefins in hydrotreated oils are formed during the hydrodeoxygenation reaction. As discussed below, any olefins in the renewable feedstock itself are preferably removed using a guard bed before the first catalyst zone. Therefore, the olefins mentioned in the oils produced in the first and second catalyst zones are intended to refer to the olefins formed, not those contained in the original feedstock.

[0022] The second catalyst zone operates under conditions favorable to hydrogenation. The effluent from the second catalyst zone contains a hydrogenated oil with an olefin content close to the thermodynamic equilibrium olefin content. Preferably, the olefin content of the hydrogenated oil has a reference olefin content in the range of 0.5 ppm-w to 5 ppm-w.

[0023] As used herein, the “reference olefin content” is calculated as the ratio of the molecular weights of thiols to octadecane. Assuming that organic sulfur-containing molecules (primarily thiols) are formed from H₂S and olefins, the olefin concentration is estimated from the organic sulfur concentration observed in the collected effluent. As an example of the calculation, this is for 105 ppm-w sulfur in the effluent. Based on one sulfur atom in octadecene (MW = 252 g / mol) and thiols (MW = 32 g / mol), the reference olefin content is calculated to be 827 ppm-w (105 * 252 / 32).

[0024] In the first catalyst zone, the hydrodeoxygenation of the renewable feedstock leads to the formation of olefins, which is driven by the recombination of sulfur with H2S present in the reactor and subsequent operations in the method, resulting in the formation of organosulfur compounds, namely thiols, disulfides, etc. The inventors have surprisingly discovered that substantially complete conversion of organic oxygen can be achieved using a first monometallic Group VIB catalyst, and that the olefins formed in the first catalyst zone can be effectively removed in the second catalyst zone of a bimetallic catalyst, leading to a more efficient hydrogenation process.

[0025] This invention advantageously avoids the undesirable formation of thiols, for example, in the catalyst bed and / or in the high-pressure separation system, through recombination with H2S. Preferably, the elemental sulfur content of the separated hydrogenated oil exiting the high-pressure separation is at most 5 ppm-w, preferably at most 2 ppm-w, corresponding to a reference olefin content of at most 40 ppm-w, preferably at most 15 ppm-w. This represents a significant advantage over conventional methods, in which the reference olefin content is greater than 800 ppm-w.

[0026] The method of the present invention is important for energy transition, as it can improve the environment by producing low-carbon energy and / or chemicals from renewable resources, particularly from biodegradable waste sources, while also improving the energy efficiency of the method.

[0027] Embodiments of the process units used to carry out the methods of the present invention are described below and / or illustrated in the accompanying drawings. For ease of discussion, additional equipment and process steps that can be used in methods for producing fuels and / or chemicals from renewable feedstocks are not shown. Additional equipment and / or process steps may include, for example, but not limited to, pre-processors, heaters, coolers, air coolers, heat exchangers, mixing chambers, valves, pumps, compressors, condensers, quench flows, recirculation flows, slip flows, purge flows, reflux flows, etc.

[0028] Figure 1 An embodiment of method 10 of the present invention is illustrated. Renewable feedstock 12 is reacted in a hydrotreating zone 14 to produce hydrotreating effluent 16. Hydrogen may be streamed and co-fed with renewable feedstock 12 before being introduced into the hydrotreating zone 14, or added to the hydrotreating zone 14 independently of renewable feedstock 12. The hydrogen may be fresh and / or recycled from another unit in the process and / or generated in an HMU (not shown). In another embodiment, hydrogen may be generated, for example, but not limited to, by water electrolysis. The water electrolysis process may be powered by renewable energy sources such as solar photovoltaic, wind, or hydropower to produce green hydrogen, nuclear energy, or by non-renewable energy sources (grey hydrogen) from other sources.

[0029] As used herein, the terms “renewable feedstock,” “renewable feed,” and “material from renewable resources” refer to feedstocks derived from renewable resources. Renewable sources can be animal, plant, microbial, and / or biologically derived or mineral-derived wastes suitable for the production of fuels, fuel components, and / or chemical feedstocks.

[0030] A preferred class of renewable materials are biorenewable fats and oils, comprising triglycerides, diglycerides, monoglycerides, free fatty acids, and / or fatty acid esters derived from biorenewable fats and oils. Examples of fatty acid esters include, but are not limited to, fatty acid methyl esters and fatty acid ethyl esters. Biorenewable fats and oils include both edible and inedible fats and oils. Examples of biorenewable fats and oils include, but are not limited to, algae oil, brown oils, low-erucic acid rapeseed oil, carinata oil, castor oil, coconut oil, rapeseed oil, corn oil, cottonseed oil, fish oil, hemp seed oil, jatropha oil, lard, flaxseed oil, milk fat, mustard oil, olive oil, palm oil, peanut oil, rapeseed oil, sewage sludge, soybean oil, soybean oil, sunflower oil, water chestnut oil, tall oil, tall oil fatty acids (TOFA), beef tallow, used cooking oils, yellow oils, white oils, and combinations thereof.

[0031] Another preferred class of renewable materials are liquids derived from biomass liquefaction and waste liquefaction processes. Examples of such liquefaction methods include, but are not limited to, (hydrogen) pyrolysis, hydrothermal liquefaction, plastic liquefaction, and combinations thereof. Renewable materials derived from biomass liquefaction and waste liquefaction processes can be used alone or in combination with biorenewable fats and oils.

[0032] In the method of the present invention, the renewable material used as a raw material may be pretreated to remove impurities, such as, but not limited to, solids, iron, chlorides, phosphorus, alkali metals, alkaline earth metals, polyethylene, and unsaponifiable compounds. Methods for removing these impurities are known to those skilled in the art.

[0033] The method of the present invention is most advantageous when processing feed streams comprising substantially 100% renewable feedstock. However, in one embodiment of the invention, the renewable feedstock may be co-processed with petroleum-derived hydrocarbons. Petroleum-derived hydrocarbons include, but are not limited to, all fractions from crude oil, natural gas condensate, tar sands, shale oil, synthetic crude oil, and combinations thereof. The present invention is more advantageous for combined renewable and petroleum-derived feedstocks comprising a renewable feed content ranging from 5% to 30% by weight. In one embodiment, the renewable feedstock is co-processed with heavy fractions from an oil refinery. For example, the petroleum-derived feedstock may be a heavy fraction from a gas oil unit.

[0034] According to the present invention, the hydrotreating zone 14 has a first catalyst zone 22 and a second catalyst zone 24. The first catalyst zone and the second catalyst zone operate in a co-current manner with a fixed-bed catalyst. Each catalyst zone 22, 24 may have a single catalyst bed or multiple catalyst beds. The hydrotreating zone 14 may be composed of a single reactor or multiple reactors. The catalyst zones 22, 24 may have different catalysts in the same catalyst bed.

[0035] In the first catalyst zone 22, hydrotreatment reaction conditions are selected to favor hydrodeoxygenation. In the second catalyst zone 24, hydrotreatment reactions are selected to favor hydrogenation. Other hydrotreatment reactions (including but not limited to hydrodenitrogenation, hydrodesulfurization, hydroaromaticization, hydrodemetallization, and combinations thereof) also occur in the hydrotreatment zone 14. Hydrodemetallization is advantageous in the first catalyst zone 22, while hydrodenitrogenation, hydrodesulfurization, and hydroaromaticization are advantageous in the second catalyst zone 24.

[0036] The catalyst in the first catalyst zone 22 is a single-metal Group VIB catalyst. As used herein, the term "single-metal" means that the active metal is from Group VIB, excluding Group VIII metals, but may include more than one Group VIB metal. Suitable Group VIB metals include Mo, W, and combinations thereof. The single-metal Group VIB catalyst may also contain a promoter metal. The single-metal catalyst may be supported or unsupported. Preferably, the single-metal catalyst is a supported catalyst. Based on the weight of the catalyst, the Group VIB metal content is in the range of 5% to 25% by weight, preferably 7% to 22% by weight, and more preferably 10% to 20% by weight.

[0037] Examples of suitable single-metal Group VIB catalysts for the first catalyst region are described in Bhan (US11,434,431B2, September 6, 2022).

[0038] The catalyst in the first catalyst region 22 can be supported on a suitable support selected from refractory oxides, molecular sieves, and combinations thereof. Examples of suitable refractory oxides include, but are not limited to, alumina, amorphous silica-alumina (ASA), titanium dioxide, silica, and combinations thereof. Examples of suitable molecular sieves include, but are not limited to, zeolite Y, zeolite β, ZSM-5, ZSM-12, ZSM-22, ZSM-23, ZSM-48, SAPO-11, SAPO-41, magnesium alkali zeolite, and combinations thereof.

[0039] The catalyst in the second catalyst zone 24 is a bimetallic Group VIII and Group VIB catalyst. As used herein, the term "bimetallic" refers to the presence of an active metal from Group VIII and an active metal from Group VIB, but may include more than one Group VIII metal and more than one Group VIB metal. The bimetallic catalyst may be supported or unsupported. Preferably, the bimetallic catalyst is a supported catalyst. Suitable Group VIII metals include Pd, Pt, Ni, Co, and combinations thereof. Preferably, the Group VIII metals are selected from Ni and Co. Suitable Group VIB metals include Mo, W, and combinations thereof.

[0040] The catalyst in the second catalyst region 24 can be supported on a suitable support selected from refractory oxides, molecular sieves, and combinations thereof. Examples of suitable refractory oxides include, but are not limited to, alumina, ASA, titanium dioxide, silica, and combinations thereof. Examples of suitable molecular sieves include, but are not limited to, zeolite Y, zeolite β, ZSM-5, ZSM-12, ZSM-22, ZSM-23, ZSM-48, SAPO-11, SAPO-41, magnesium alkali zeolite, and combinations thereof.

[0041] Compared to their oxide forms, monometallic and bimetallic hydrotreating catalysts in sulfidated form are generally more active. The catalyst metal is typically in an oxide state when loaded into the reactor and is preferably activated by reducing or sulfiding the metal oxide. The sulfidation process is used to convert the catalyst from a calcined oxide state to an active sulfidated state. The catalyst can be pre-sulfided or sulfided in situ. Because renewable feedstocks typically have low sulfur content, sulfiding agents are often added to the feed to maintain the catalyst in its sulfidated form.

[0042] Hydrotreating catalysts can be sulfided in situ or ex-situ. In-situ sulfidation can be achieved by supplying a sulfur source to the hydrotreating catalyst during process operation, typically H2S or H2S precursors (i.e., compounds that readily decompose into H2S, such as, for example, dimethyl disulfide, di-tert-nonyl polysulfide, or di-tert-butyl polysulfide). The sulfur source can be supplied along with the feed, the hydrogen stream, or separately. Another suitable sulfur source is a sulfur-containing hydrocarbon stream co-fed with the feedstock, whose boiling point falls within the boiling range of diesel or kerosene. Furthermore, the addition of sulfur compounds to the feedstock helps control catalyst stability and reduces hydrogen consumption.

[0043] Preferably, H2S is supplied to the reactor in an amount ranging from 50 ppm-v to 5,000 ppm-v, more preferably from 100 ppm-v to 3,000 ppm-v, and more preferably from 500 ppm-v to 2,000 ppm-v. The amount of H2S supplied depends on many factors, including, but not limited to, the type and amount of catalyst metal in the hydrotreating step, the operating temperature, and other operating conditions.

[0044] Preferably, the volume of catalyst in the first catalyst zone 22 is greater than the volume of catalyst in the second catalyst zone 24. More preferably, the volume ratio of the single-metal catalyst to the bimetallic catalyst is in the range of 60:40 to 95:5, and most preferably in the range of 70:30 to 90:10. The volume ratios referred to herein refer to the volumes of the supported single-metal catalyst and the bimetallic catalyst. The volume ratios do not include any volume of any guard bed catalyst or non-catalytic trapping material.

[0045] In one embodiment, the hydrotreating zone 14 further includes a guard bed prior to the first catalyst zone to demetallize the renewable feedstock and / or hydrogenate any bulk olefins in the renewable feedstock. The guard bed may be a low-activity catalyst and / or a non-catalytic trap. The active metal for hydrogenation includes catalytically active metals of Group VIII and / or Group VIB, including but not limited to Ni, Co, Mo, W, and combinations thereof. The metal content in the guard bed is typically low, for example, 2.5% to 5% by weight, and a support with high porosity is typically used to trap the captured metal from the renewable feedstock. Suitable supports include refractory oxides, molecular sieves, and combinations thereof. Examples of suitable refractory oxides include, but are not limited to, alumina, ASA, titanium dioxide, silica, and combinations thereof. Examples of suitable molecular sieves include, but are not limited to, zeolite Y, zeolite β, ZSM-5, ZSM-12, ZSM-22, ZSM-23, ZSM-48, SAPO-11, SAPO-41, magnesium alkali zeolite, and combinations thereof.

[0046] Operating conditions in the hydrogenation reactor include pressures ranging from 1.0 MPaG to 20 MPaG, temperatures ranging from 200°C to 410°C, and 0.3 m³ / h based on fresh feed. 3 / m 3 .h to 5m 3 / m 3 The liquid hourly space velocity is in the range of h. Preferably, the pressure is selected from the range of 2.0 MPaG to 15 MPaG. Preferably, the temperature is in the range of 200°C to 400°C.

[0047] The ratio of hydrogen supplied to feed in the hydrotreating zone 14 is in the range of 200 standard L to 10,000 standard L (under standard conditions of 0°C and 1 atm (0.101 MPa)) / kg feed. Feed referred to herein is the total amount of fresh feedstock, excluding any diluents that may be added. Preferably, hydrogen is added at the top of the hydrotreating zone 14.

[0048] As described above, the method of the present invention is particularly advantageous for methods that require a separation zone prior to subsequent processes. For example, but not limited to, in cases where the hydrotreated effluent is passed to a hydroisomerization zone with a noble metal catalyst, the hydrotreated effluent is typically passed to a separation zone prior to the hydroisomerization zone. In conventional methods where H2S is present in the hydrotreated effluent, problems arise in the separation zone when sulfur recombines with olefins to form harmful thiols.

[0049] Therefore, in a preferred embodiment of the invention, the hydrotreated oil from the second catalyst zone 24 is directed to the separation zone 30 for separating the products of the hydrotreatment reaction into a gaseous effluent and a liquid hydrotreated effluent 16. When the catalyst used for hydroisomerization contains a noble metal, the separation zone 30 is provided to remove or at least significantly reduce components that poison or otherwise adversely affect the hydroisomerization catalyst. In cases where non-noble metals are used for hydroisomerization, the separation zone 30 is optional. When the hydrotreated oil is passed to a thermal cracker (such as a steam cracker) to produce olefins, it is desirable to separate a light fraction from the liquid effluent in the separation zone 30.

[0050] When separation zone 30 is included, separation zone 30 has one or more separation units, including, for example but not limited to, gas / liquid separators, including hot high-pressure and low-pressure separators, medium-pressure and low-pressure separators, cold high-pressure and low-pressure separators, high-pressure and low-pressure strippers, integrated strippers, and combinations thereof. Integrated strippers include strippers integrated with hot high-pressure and low-pressure separators, medium-pressure and low-pressure separators, and cold high-pressure and low-pressure separators. Those skilled in the art will understand that the high-pressure separator operates at a pressure close to that of the hydrotreating zone 14, suitably 0 to 10 bar (0 MPa to 1 MPa) lower than the reactor outlet pressure, while the low-pressure separator operates at a pressure lower than that of the preceding reactor or the preceding high-pressure separator in the hydrotreating zone 14, suitably 0 to 15 bar (0 MPaG to 1.5 MPaG). Similarly, those skilled in the art will understand that "hot" means that the hot separator operates at a temperature close to that of the preceding reactor in the hydrotreating zone 14, suitably sufficiently above the water dew point (e.g., ≥10°C, preferably ≥20°C) and sufficiently above the salt deposition temperature (e.g., ≥10°C, preferably ≥20°C), while the intermediate and cold separators are at reduced temperatures relative to the preceding reactor in the hydrotreating zone 14. For example, the cold separator is suitably at a temperature achievable via an air cooler. "Intermediate temperature" should be understood to refer to any temperature between the hot and cold separator temperatures.

[0051] Additionally, separation zone 30 may include one or more processing units, including, but not limited to, membrane separation units, amine scrubbers, pressure swing adsorption (PSA) units, alkaline scrubbing, and combinations thereof. The processing units are preferably selected to separate the desired gaseous molecules. For example, an amine scrubber is used to selectively separate H2S and / or carbon oxides from H2 and / or hydrocarbons. As another example, a PSA unit may be used to purify the hydrogen stream for recycling to the stripper and / or reactor in hydrotreating zone 14.

[0052] A portion of the hydrotreatment effluent 16 from one or more separator units may be returned to the hydrotreatment zone 14, for example, as a quench stream (not shown) or as a diluent (not shown) for feedstock 12. The recirculation and quench streams can be prepared in a manner known to those skilled in the art. Preferably, the recirculation stream is combined with a renewable feedstock. Alternatively, the recirculation stream may be distributed and added to the hydrotreatment zone 14 at two or more inlets. One or more quench streams may be added to the catalyst beds in the first catalyst zone 22 and / or the second catalyst zone 24 in a manner known to those skilled in the art.

[0053] The hydrotreated effluent 16 (with or without a separation step) is transferred to the hydroisomerization zone 34 under hydroisomerization conditions that induce the hydroisomerization reaction. The hydroisomerization reaction increases the branching of alkane compounds produced by the hydrotreated zone 14, thereby improving the cold flow properties of the fuel.

[0054] The hydroisomerization catalyst can be any suitable catalyst composition known to those skilled in the art. Preferably, the hydroisomerization catalyst comprises a Group VIII metal. More preferably, the hydroisomerization catalyst further comprises a zeolite material. The hydroisomerization catalyst may further comprise a binder and / or a support, such as, but not limited to, silica, alumina, silica-alumina, and combinations thereof. Preferably, the Group VIII metal is selected from the group consisting of platinum, palladium, nickel, and combinations thereof. When the Group VIII metal is Ni, the hydroisomerization preferably includes a Group VIB metal, preferably Mo or W.

[0055] The zeolite material is preferably selected from the group consisting of the following: β, COK-7, EU-1, EU-2, EU-11, IZM-1, MCM-22, NU-10, ZSM-5, ZSM-12, ZSM-22, ZSM-23, ZSM-30, ZSM-35, ZSM-48, ZSM-50, ZSM-57, and combinations thereof.

[0056] The catalyst can be the same or different throughout the hydroisomerization zone 34. The hydroisomerization zone 34 can include a single catalyst bed or multiple catalyst beds. The catalyst can be the same throughout a single catalyst bed, optionally a mixture of catalysts may be present, or different catalysts may be provided in two or more layers within the catalyst bed. In embodiments with multiple catalyst beds, the catalyst in each catalyst bed can be the same or different.

[0057] The hydroisomerization zone 34 operates in the presence of hydrogen at pressures ranging from 1 MPaG to 30 MPaG and temperatures ranging from 260°C to 400°C. Preferably, the pressure is in the range of 2 MPaG to 17 MPaG and the temperature is in the range of 300°C to 380°C. Based on effluent 16, LHSV is generated in 0.2 h. -1 up to 4h -1 Within the range. The ratio of hydrogen to liquid supplied to the hydroisomerization zone 34 is in the range of 100 standard L to 1500 standard L (under standard conditions of 0°C and 1 atm (0.1 MPa)) / kg liquid.

[0058] Hydroisomerization is particularly beneficial for improving the production of kerosene for jet fuel. To increase kerosene yield, the WABT (weighted average bed temperature) in the hydroisomerization zone is typically increased. WABT is the representative temperature of the catalyst bed assuming no losses or gains from the surrounding environment in an adiabatic reactor. Those skilled in the art will understand that the temperature distribution typically increases from inlet to outlet.

[0059] The product from the hydroisomerization region 34 is directed to the post-processing section. Various embodiments of the post-processing section may be considered. For example, but not limited to, the post-processing section may be as described in WO2023 / 043792 or WO2023 / 043764, published on March 23, 2023, which are incorporated herein by reference.

[0060] The post-processing section includes one or more product recovery zones 36, thereby generating the desired product stream. For example, Figure 1 The embodiment illustrates an exhaust gas stream 42, a naphtha boiling point range stream 44, a kerosene boiling point range stream 46, a diesel boiling point range stream 48, and a heavy fraction 52. In this embodiment, the exhaust gas stream 42 suitably contains C1-C5 hydrocarbons, while the naphtha boiling point range stream 44 suitably contains C4-C5 hydrocarbons with a boiling point range of -12°C to 204°C. 12 Hydrocarbons. Kerosene with a boiling point range of 46 is preferably composed of C6-C hydrocarbons with a boiling point range of 90°C to 300°C. 18 Hydrocarbon composition. In one embodiment, the diesel boiling point range stream comprises C8-C hydrocarbons with a boiling point range of 120°C to 400°C. 26Hydrocarbons. In this embodiment, heavy fraction 52 has a C60 concentration with a boiling point greater than 250°C. 17 + hydrocarbon.

[0061] exist Figure 2 In the implementation scheme, the heavy fraction 52 is recycled to the hydroisomerization zone 34. Figure 3 and Figure 4 In the implementation scheme, the heavy fraction is recycled to the hydrocracking zone 54 before contacting the hydrotreated effluent 16.

[0062] In a preferred embodiment, method 10 relates to increasing the yield of kerosene boiling point range stream 46. Specifically, in a preferred embodiment, method 10 relates to a kerosene product conforming to ASTM D7566, wherein the alkane kerosene synthesized from hydrogenated esters and fatty acids has a maximum T10 distillation temperature of 205°C (using ASTM Test Method D86) and a maximum final boiling point of 300°C.

[0063] In this case, such as Figure 2 and Figure 3 As shown, at least a portion of the diesel boiling range stream 48 is recycled together with the heavy fraction 52. Another portion of the diesel boiling range stream 48 may be discharged as a discharge stream. The product recovery zone 36 may include further separating the diesel boiling range stream 48 into a light diesel stream, which may be discharged, for example, as a discharge stream, while the heavy diesel stream is recycled together with the heavy fraction 52.

[0064] Alternatively, product recovery zone 36 may include further separation of heavy contaminants from heavy fraction 52. Depending on the feedstock and / or processing conditions, there may be heavy contaminants that may not be reactive in subsequent hydrocracking and / or hydroisomerization zones. In this case, it is preferable to provide an effluent stream of the heaviest portion of heavy fraction 52. In one embodiment, the effluent stream has substantially the same composition as heavy fraction 52. In another embodiment, the effluent stream may be a product of further processing and / or separation of heavy fraction 52 to selectively remove contaminants from it.

[0065] In one implementation scheme, such as Figure 4 As shown, hydrocarbons in the diesel boiling point range can be part of the heavy fraction 52 and are recycled for cracking and isomerization to disappear.

[0066] Figures 2 to 4 The product stream and recycling options shown are examples and can be implemented in... Figures 2 to 4 Other implementations may modify and / or use this. For example, such as Figure 4 The option to recycle hydrocarbons within the diesel boiling point range shown is available. Figure 2 and Figure 3 This is implemented in the proposed solution. Alternatively, such as... Figure 2 and Figure 3 The diesel boiling point range shown can be flowed with or without a discharge flow from... Figure 4 The product is discharged into the product recovery area specified in the implementation plan.

[0067] As described above, embodiments of product recovery zone 36 may consist of one or more unit operations. For example, product recovery zone 36 may include a product stripper for stripping entrained and / or dissolved gases from the hydroisomerization zone effluent, a naphtha stripper for generating stripper exhaust gas and naphtha stream, a naphtha stabilization column, a naphtha rectification column, a naphtha recovery column, a column top separator, a vacuum fractionator, an atmospheric fractionator, and combinations thereof.

[0068] exist Figure 3 and Figure 4 In one embodiment, at least a portion of the heavy fraction 52 is introduced into a hydrocracking zone 54 under hydrocracking conditions sufficient to induce a hydrocracking reaction to produce a hydrocracking effluent. In another embodiment, a further portion of the heavy fraction 52 may be directed to further processing for increasing the value of the heavy fraction 52.

[0069] The hydrocracking catalyst can be any suitable catalyst composition known to those skilled in the art. Preferably, the hydrocracking catalyst comprises a Group VIII metal. More preferably, the hydrocracking catalyst further comprises an acidic material.

[0070] The acidic material can be an amorphous acidic material, a crystalline acidic material, or a combination thereof. The amorphous acidic material can be, for example, but not limited to, ASA. The crystalline acidic material can be selected from the group consisting of: β, COK-7, EU-1, EU-2, EU-11, IZM-1, MCM-22, NU-10, ZSM-5, ZSM-12, ZSM-22, ZSM-23, ZSM-30, ZSM-35, ZSM-48, ZSM-50, ZSM-57, and combinations thereof.

[0071] Preferably, the Group VIII metal is selected from the group consisting of platinum, palladium, nickel, and combinations thereof. When the Group VIII metal is Ni, the hydroisomerization catalyst preferably contains a Group VIB metal, preferably Mo or W.

[0072] The hydrocracking catalyst may further comprise a binder and / or support, such as, but not limited to, silica, alumina, silica-alumina, and combinations thereof.

[0073] The hydrocracking zone 54 operates in the presence of hydrogen at a pressure ranging from 1 MPaG to 30 MPaG and a temperature ranging from 260°C to 400°C. Preferably, the pressure is in the range of 2 MPaG to 18 MPaG and the temperature is in the range of 280°C to 400°C.

[0074] Hydrocracking conditions and catalysts are selected to favor cracking rather than branching.

[0075] exist Figure 3 In one embodiment, the hydrocracking zone 54 is configured as a single-stage reactor above the hydroisomerization zone 34. In this embodiment, the heavy fraction 52 is recycled to a single-stage reactor comprising both the hydrocracking zone 54 and the hydroisomerization zone 34. The heavy fraction 52 is first reacted in the hydrocracking zone 54. Then, the hydrocracking effluent from the hydrocracking zone 54 is passed to the hydroisomerization zone 34, where the hydrocracking effluent is combined with the hydrotreated effluent 16 to simultaneously isomerize both the hydrocracking effluent and the hydrotreated effluent 16.

[0076] exist Figure 4 In this embodiment, the hydrocracking zone 54 and the hydroisomerization zone 34 are provided in a two-stage configuration. In this embodiment, the heavy fraction 52 is first reacted in the hydrocracking zone 54. Then, the hydrocracking effluent from the hydrocracking zone 54 is passed to the hydroisomerization zone 34, where the hydrocracking effluent is combined with the hydrotreated effluent 16 to isomerize the hydrocracking effluent and the hydrotreated effluent 16.

[0077] Hydroisomerization zone 34 optionally includes a hydrorefining zone (not shown). During the hydroisomerization step and / or depending on the feedstock used (e.g., cashew oil), some aromatics and / or trace amounts of olefins may be present in the effluent of the hydroisomerization zone. In this case, a hydrorefining step is preferably provided to reduce the aromatic content of the product stream.

[0078] The hydrorefining components can be used in bulk metallic form or the metal can be supported on a support. The active metals for hydrogenation include catalytically active metals of Group VIII and / or Group VIB, including but not limited to Ni, Co, Mo, W, and combinations thereof. Preferably, Group VIII metals are selected from the group consisting of platinum, palladium, nickel, and combinations thereof. Suitable supports include refractory oxides. Examples of suitable refractory oxides include, but are not limited to, alumina, ASA, titanium dioxide, silicon dioxide, and combinations thereof.

[0079] Another embodiment of the present invention is in Figure 5As illustrated, the hydrotreated effluent 16 is passed to a steam cracker 64 to produce olefins. Optionally, a product stream of ultra-low sulfur diesel 62 is produced, wherein the residual stream is set to the steam cracker 64.

[0080] The hydrotreated effluent 16 and dilution gas (preferably steam) are preheated to a temperature of 140°C to 375°C at a pressure ranging from 1 MPaG to 2 MPaG to produce a gas / liquid mixture. The preheated gas / liquid mixture is separated in a gas-liquid separator. The separated gas is fed into the convection zone of a steam cracker 64, and then into the radiant zone of the steam cracker 64, where the gas is pyrolyzed into products containing olefins. Example

[0081] For illustrative purposes only, the following non-limiting embodiments of the method of the invention protected by the claims are provided. Comparative Example 1

[0082] In the reactor, a single-metal hydrotreating catalyst was placed. The catalyst bed consisted of 93.75 mL of hydrotreating catalyst containing 14.5 wt% Mo on ASA / alumina. The hydrotreating catalyst bed was diluted with 0.2 mm silicon carbide particles at a ratio of 1:1.5. The silicon carbide particles were applied to mitigate the reactor wall effect, which can disrupt the uniform liquid distribution across the cross-section of the catalyst bed.

[0083] The temperature of the catalyst bed is controlled by an oven to operate at 300°C.

[0084] A feedstock consisting of refined tallow was supplied to the catalyst bed at a WHSV ratio of 1.33 g fresh oil / mL catalyst / hour. The feedstock was doped with 0.34 wt% SULFRZOL as a hydrogen sulfide precursor. ® To maintain catalyst sulfidation, the sulfur concentration in tallow is 3.6 ppm-w. The reactor is operated by recycling the liquid fraction of the hydrocarbon effluent at a 1:1 volume ratio compared to the fresh feed. Based on the fresh feed, a gas stream containing 100% hydrogen by volume is supplied at a gas-oil ratio of 875 NL / kg. The total pressure at the reactor outlet is 75 bar (gauge pressure).

[0085] The degree of conversion of beef tallow feedstock was determined using multiple methods. Pyrolysis and UV fluorescence analysis of the hydrocarbon liquid were used to determine the concentrations of organic oxygen and sulfur. Gas chromatography was used to analyze the gaseous effluent.

[0086] The conversion rate of organic oxygen was greater than 99%, with an elemental oxygen concentration of 900 ppm-w and an organic sulfur concentration of 105 ppm-w in the effluent. The selectivity for carbon oxide formation, as determined by gas chromatography analysis of the gaseous effluent, was 3%, indicating that almost all C... 16 and C 18 Hydrocarbons were retained. The nature of the sulfur components was determined using two-dimensional gas chromatography with sulfur-sensitive detection to confirm that the sulfur in the effluent was not in the form of hydrogen sulfide or unconverted sulfur. ® It exists in the form of. Example 2

[0087] In one embodiment of the invention, Comparative Example 1 was repeated using a different catalyst pack. In this case, a hydrogenation catalyst bed consisting of 68.75 mL of a hydrogenation catalyst containing 14.5 wt% Mo on ASA / alumina was placed above a bed of 25 mL of a conventional hydrotreating catalyst containing 4 wt% Ni and 15 wt% Mo on alumina. The catalyst in the hydrodeoxygenation bed was diluted 1:1.5 with 0.2 mm diameter silicon carbide particles, while the catalyst in the hydrogenation catalyst bed was diluted 1:1.5 with 0.2 mm diameter silicon carbide particles.

[0088] The reactor operates at a temperature of 300°C.

[0089] In this embodiment, the organic oxygen conversion rate is again greater than 99%, wherein the elemental oxygen concentration in the liquid effluent is less than the detection limit of 300 ppm-w, and the organic sulfur concentration is 1 ppm-w. The selectivity for the formation of carbon oxides is 12%.

[0090] Surprisingly, replacing part of the Mo on the ASA / alumina catalyst with a conventional hydrotreating catalyst that has stronger hydrogenation capabilities improves the robustness of the process by eliminating trace levels of olefins that are typically present in the effluent of hydrotreating processes and thus preventing hydrogen sulfide from recombinating with these olefins into organosulfur compounds such as thiols.

[0091] While various embodiments and uses have been described with reference to them, it should be understood that these embodiments are exemplary and the scope of the invention is not limited thereto. Many variations, modifications, additions, and improvements are possible. Various combinations of the techniques provided herein can be used.

Claims

1. A method for hydrogenating renewable feedstock, the method comprising the following steps: Provide renewable raw materials; The renewable feedstock is passed to a first catalyst zone containing a single-metal Group VIB catalyst under hydrodeoxygenation conditions to produce a first reaction mixture containing a hydrotreated oil with an olefin content at least 250 times greater than the thermodynamic equilibrium olefin content; and The first reaction mixture is passed to a second catalyst zone containing a bimetallic catalyst having Group VIII and Group VIB metals under hydrogenation conditions to produce an effluent containing a hydrogenated oil with an olefin content close to thermodynamic equilibrium.

2. The method according to claim 1, wherein the olefin content of the hydrotreated oil exceeds the thermodynamic equilibrium olefin content by at least 500 times.

3. The method according to claim 1, wherein the hydrotreated oil has a reference olefin content of at least 125 ppm-w.

4. The method of claim 1, wherein the hydrogenated oil has a reference olefin content in the range of 0.5 ppm-w to 5 ppm-w.

5. The method according to claim 1, wherein the volume of the monometallic catalyst is greater than the volume of the bimetallic catalyst.

6. The method according to claim 1, wherein the volume ratio of the monometallic catalyst to the bimetallic catalyst is in the range of 60:40 to 95:5, preferably in the range of 70:30 to 90:

10.

7. The method of claim 1, further comprising the step of separating the hydrogenated oil from the gas phase present in the effluent from the second catalyst zone.

8. The method of claim 7, wherein the separation step includes a high-pressure separator.

9. The method of claim 7, wherein the thiol content of the separated oil has a reference olefin content of up to 40 ppm-w, preferably up to 15 ppm-w.

10. The method of claim 1, wherein the renewable raw material is selected from the group consisting of one or more biorenewable fats and oils, liquids derived from biomass liquefaction processes, liquids derived from waste liquefaction processes, and combinations thereof.

11. The method of claim 1, further comprising, preferably, adding a petroleum-derived feedstock for co-processing with the renewable feedstock in an amount that produces a feed stream comprising 5% to 30% by weight of the renewable feedstock.

12. The method of claim 7, further comprising the step of transferring the hydrogenated oil to the hydroisomerization step.

13. The method of claim 7, further comprising the step of transferring the hydrogenated oil to a steam cracking step.

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