Catalyst and process for the catalytic conversion of naphtha to light olefins

Abstract

A zeolite catalyst and methods of use thereof are provided, wherein the metal, preferably platinum, on the catalyst is not well dispersed as shown by low H2 chemisorption levels and low chemisorbed hydrogen to platinum ratios. In a process for converting naphthas to light olefins, the catalyst produces a higher proportion of ethane, the process comprising contacting a naphtha stream with a zeolite catalyst under conditions to dehydrogenate the naphtha to olefins, interconvert the olefins to lighter olefins, and hydrogenate the lighter olefins to produce a light paraffin stream comprising ethane and propane, thereby producing a light paraffin stream.

Description

[0001] Priority Statement

[0002] This application claims the benefit of U.S. Provisional Patent Application Serial No. 63 / 616,650, filed on December 31, 2023, the entire disclosure of which is incorporated herein by reference. Technical Field

[0003] This field relates to the conversion of naphtha to light olefins. Particularly relevant to this field is the conversion of naphtha to light olefins via a two-step conversion process. More specifically, this field relates to the use of catalysts that maximize ethylene production. Background Technology

[0004] Light olefin production is crucial for producing enough plastics to meet global demand. Alkane dehydrogenation (PDH) is a method in which light alkanes (such as propane and butane) can be dehydrogenated to produce propylene and butene, respectively. Dehydrogenation is an endothermic reaction that requires external heat to drive the reaction.

[0005] Fluid catalytic cracking (FCC) is another endothermic method that can be fine-tuned to produce large quantities of propylene. However, not every FCC unit can be fine-tuned to produce large amounts of propylene. Moreover, high-propylene FCC units do not produce much ethylene; less than 1% of the global ethylene supply comes from FCC.

[0006] A significant amount of ethylene consumed in the production of plastics and petrochemicals such as polyethylene is produced through the thermal cracking of hydrocarbons. Steam is typically mixed with the feed stream into the cracker to reduce hydrocarbon partial pressures and increase olefin yields, while minimizing the formation and deposition of carbonaceous material in the cracking reactor. Therefore, this method is often referred to as steam cracking or pyrolysis. Ethane oxidative dehydrogenation is a newer catalytic method for converting ethane to ethylene, which can be performed at lower temperatures with significantly reduced carbon oxide emissions compared to steam cracking.

[0007] Two types of feedstocks are commonly used in steam cracking. Ethane feedstocks are used in regions where light hydrocarbon gases are abundant. In regions where natural gas is scarce, naphtha feedstocks are used for steam cracking. Pyrolytic naphtha cracking has long determined the price of ethylene in the industry due to its higher production costs compared to pyrolytic ethane cracking. Currently, the world does not produce enough ethane to meet the growing demand for ethylene. Therefore, regions with insufficient ethane supply, such as Asia and Europe, rely primarily on naphtha cracking to supply ethylene. Naphtha cracking produces only 30% to 35% ethylene; the remainder consists of both high-value byproducts including propylene, butadiene, and butene-1, and relatively low-value byproducts including pyrolysis oil, pyrolysis gas, and fuel gas. Additional pressures on naphtha cracking, including minimum production requirements and environmental concerns, have led to government refusal to approve it in some regions, such as China. The ethylene industry needs a more efficient, economical, and environmentally friendly route to light olefins from naphtha feedstocks.

[0008] Naphtha to ethane and propane (NEP) is a highly efficient proprietary ethylene and propylene production pathway when combined with ethane steam cracking and propane dehydrogenation, resulting in significantly improved molecular conversion efficiency. To achieve primary ethane production (high ethane and propane) and meet the greater demand for ethylene (compared to propylene), maximizing ethylene yield, NEP is needed to produce high ethane and low propane. This disclosure provides a catalyst with the property of producing high levels of ethane and propane and low amounts of methane. Summary of the Invention

[0009] A method for converting naphtha into light olefins, the method comprising: contacting a naphtha stream with a catalyst under conditions in which the alkanes are dehydrogenated to olefins, the olefins are interconverted to lighter olefins, and the lighter olefins are hydrogenated to produce a light alkane stream comprising ethane and propane, thereby producing a light alkane stream. It has been found that catalysts with the following characteristics can be used to produce larger quantities of ethylene compared to other catalysts. Generally, it is assumed that when the metal is well dispersed, the metal deposited on the catalyst provides a higher yield or greater selectivity. However, in this disclosure, it has been found that less dispersed metal produces a higher ethane yield as desired herein. Dispersion is represented by the amount of hydrogen adsorbed by the catalyst; when the metal is not well dispersed, the amount of adsorbed hydrogen is lower. The catalyst contains greater than 0.02 wt% platinum and has an acidity greater than 0.15 mmol / g. The catalyst has a low chemisorption ratio of hydrogen to platinum, and preferably has a metal characteristic of a chemisorption ratio of less than 0.45 to platinum. The catalyst is further described as having metallic properties with less than 4.5 μmol / g of chemisorbed hydrogen. Attached Figure Description

[0010] Figure 1 This is a schematic diagram of the method and apparatus disclosed herein.

[0011] Figure 2 yes Figure 1 A schematic diagram of the alternative implementation method and equipment.

[0012] Figure 3 yes Figure 1 A schematic diagram of the method and apparatus for an additional alternative implementation scheme.

[0013] Figure 4 yes Figure 1 A schematic diagram of the alternative implementation method and equipment.

[0014] definition

[0015] The term "connectivity" refers to the operative permission for fluid flow between enumerated components, which can be characterized as "fluid connectivity".

[0016] The term "downstream connectivity" means that in downstream connectivity, at least a portion of the fluid flowing toward the body can be operatively flowed from the object with which it is fluidly connected.

[0017] The term "upstream connectivity" means that at least a portion of the fluid flowing out of the main body can be operatively directed to an object in fluid communication with it.

[0018] The term "direct connection" means that fluid flow from an upstream component enters a downstream component without passing through any other intermediary container.

[0019] The term "indirect connection" refers to fluid flow from an upstream component entering a downstream component after passing through an intermediary container.

[0020] The term "bypass" means that an object is disconnected from the downstream entity at least within the scope of the bypass.

[0021] As used herein, the terms “major” or “most” mean greater than 50%, appropriately greater than 75%, and preferably greater than 90%.

[0022] The term "Cx" should be understood as referring to a molecule having the number of carbon atoms indicated by the subscript "x". Similarly, the term "Cx-" refers to a molecule containing less than or equal to x, and preferably x and fewer carbon atoms. The term "Cx+" refers to a molecule having more than or equal to x, and preferably x and more carbon atoms.

[0023] The term "tower" refers to one or more distillation columns used to separate one or more components with different volatility. Unless otherwise specified, each column includes a condenser at the top of the column for condensing a portion of the overhead stream and refluxing it back to the top of the column, and a reboiler at the bottom of the column for vaporizing a portion of the bottom stream and returning it to the bottom of the column. The feed to the column can be preheated. The top pressure is the pressure of the vapor at the top of the column at the vapor outlet. The bottom temperature is the liquid temperature at the bottom outlet. Top and bottom lines refer to the net lines from any downstream reflux or reboiler to the column. A stripping column may omit the reboiler at the bottom of the column, instead providing the heating requirements and separation power for a liquefied inert medium such as steam. Stripping columns typically feed from the top tray and remove the main product from the bottom.

[0024] As used herein, the term "separator" refers to a vessel having an inlet and at least one top vapor outlet and a bottom liquid outlet, and may also have an outlet for a water stream from a boot. A flash tank is a type of separator that can be connected downstream to a separator that can operate at higher pressures. Detailed Implementation

[0025] In the proposed method, C3-C8+ hydrocarbon feedstock is first loaded into a naphtha-to-ethane and propane unit to convert naphtha into the desired ethane and propane, and less desired methane. The resulting ethane is then fed into an ethylene production unit. These units provide an ethane-to-ethylene yield of over 75%. The resulting propane is fed into a propylene production unit, which provides a propane-to-propylene yield of over 85%. Methane byproducts from the naphtha conversion unit and the ethane and propane production units can be used as fuel, including fuel required for the operation of ethylene and propylene production units at high temperatures. Unconverted or underconverted C4+ components at the reactor outlet can be recovered to the reactor inlet for further processing into ethane and propane. Aromatics can also be recovered and further processed.

[0026] In the method proposed herein, it is preferable to contact the feed stream with a catalyst having low platinum dispersion to achieve low methane yield at a given ethane or ethane / propane yield level when tested under metal-functional-constrained conditions (low hydrogen / hydrocarbon ratio). Low platinum dispersion can be achieved by steaming the catalyst after platinum incorporation (if using a Pt(2+) precursor to exchange platinum ions onto the zeolite micropores). Steaming the platinum-incorporated catalyst drives platinum, initially located inside the zeolite, to the outside of the zeolite crystals and potentially to the binder. Alternatively, platinum can be introduced onto the binder using a platinum precursor that selectively chemisorbs onto the binder (such as chloroplatinic acid in silica-bonded zeolite catalysts). Furthermore, by incorporating relevant competing ions, such as NH4NO3 (for Pt(II) precursors) and HCl (for CPA precursors), the platinum distribution between the zeolite micropores and the zeolite exterior / binder can be tuned to minimize methane yield while maximizing ethane yield. When operating naphtha to ethane and propane reactions under acidity-limited conditions (high H2 / hydrocarbon ratio), it is desirable to operate with a catalyst having a minimum acidity of 0.10 mmol / g, preferably 0.12 mmol / g, and most preferably 0.15 mmol / g, as measured by NH3-TPD. Operating with a low-acidity catalyst would require high temperatures, resulting in high methane production at the target ethane yield.

[0027] Go to Figure 1 The naphtha stream in line 10 can be combined with the hydrogen stream in line 22 and the heavy stream in line 12 to provide the charge stream in line 11, which is heated and fed into the naphtha to ethane and propane (NEP) reactor 16 for contact with the NEP catalyst. The naphtha stream may include C4 to C12 hydrocarbons, preferably having a T10 between -10°C and 60°C and a T90 between 70°C and 180°C. The naphtha feed stream may contain n-alkanes, isoalkanes, cycloalkanes, and aromatics. The naphtha stream can be heated to a reaction temperature of 300°C to 600°C, suitably between 325°C and 550°C, and preferably between 350°C and 525°C. The weight hourly space velocity (WHSV) should be between 0.3 hr⁻¹ and 20 hr⁻¹, suitably between 0.5 hr⁻¹ and 10 hr⁻¹, and preferably between 1 hr⁻¹ and 4 hr⁻¹. The total pressure should be between 0.1 MPa and 3 MPa (absolute pressure), preferably not exceeding 2 MPa (absolute pressure). Under these conditions, the C2-C4 yield is always greater than 80% by weight, while the methane yield is less than 10% by weight, suitably less than 8% by weight, and typically less than 6% by weight, and preferably not exceeding 5% by weight. Under these conditions, ethane may constitute more than 60% by weight of the total C2 to C3 and of the total C2 to C4 produced in NEP reactor 16.

[0028] The molar ratio of hydrogen to hydrocarbons is important for the production of ethane and propane. The hydrogen-to-hydrocarbon ratio should be from 0.3 to 15, and preferably from 0.5 to 5. In another embodiment, the molar ratio of hydrogen to hydrocarbons typically does not exceed 5, suitably does not exceed 3, and preferably does not exceed 2. A low hydrogen-to-hydrocarbon ratio promotes the desired reaction kinetics initiated by dehydrogenation. The hydrogen-to-hydrocarbon ratio can be in the range of 50% to 500% of the stoichiometric requirements for converting naphtha molecules to ethane and / or propane, suitably not exceeding 300%, and preferably not exceeding 200%.

[0029] The molar ratio of hydrogen to hydrocarbons depends on the feed type, including alkanes, cycloalkanes, or aromatics, the carbon number of the feed molecules, and the desired products between ethane-dominant, propane-dominant, or ethane and propane of comparable abundance, as illustrated in Table 1 below. For example, to stoichiometrically convert 1 mole of propane to ethane, this method would require a co-feed of 0.5 moles of hydrogen. In practice, the method can operate above or below this stoichiometry of 0.5, such as 0.33, to obtain greater than 40% ethane and less than 15% methane, depending on process design parameters such as feed contaminants, reactor type (fixed bed, moving bed, fluidized bed), and regeneration frequency. As the carbon number of the feed molecules increases from light naphtha (C5-C7) to the full range of naphtha (C6-C7), the yield increases accordingly. 10 The amount of hydrogen required for the reaction increases. For example, completely converting 1 mole of nonane to ethane and propane will require 3.5 moles of hydrogen and 2.0 moles of hydrogen, respectively. The disclosed method can be operated at three to five times the hydrogen-to-hydrocarbon ratio required to convert the feed molecules stoichiometrically to ethane and propane, respectively. The method can also be operated at 50% of the hydrogen-to-hydrocarbon ratio required to convert the feed molecules stoichiometrically to ethane and propane, respectively. The hydrogen-to-hydrocarbon ratio will also depend on the need for producing petrochemical aromatics such as benzene, toluene, and xylene. Following these guidelines, the disclosed method can be used to convert light naphtha (C5-C7) and full-range naphtha (C5-C6) into hydrocarbons. 10 The light and heavy naphtha derived from hydrocracking are converted into ethane, propane, petrochemical aromatics and mixtures thereof, the light and heavy naphtha including alkane molecules, cycloalkanes molecules and aromatic molecules.

[0030] Table 1

[0031]

[0032] The NEP catalyst dehydrogenates alkane molecules in naphtha to their olefin analogs, interconverts olefins to lighter olefins and hydrogenates the lighter olefins to produce a light alkane stream containing ethane and propane. Interconversion may mean that olefins are also oligomerized to higher-carbon olefins, and then these higher-carbon olefins are cracked to lower-carbon olefins. This chemical mechanism avoids or minimizes hydrocracking and pyrolysis reactions that produce methane. Methane is an undesirable byproduct that represents an opportunity to produce valuable ethane and propane and consumes excess hydrogen. The NEP catalyst can also dehydrogenate naphthenes to aromatics such as benzene, toluene, and xylene.

[0033] The NEP catalyst for converting naphtha to ethane and propane may comprise a molecular sieve, including a macroporous or mesoporous molecular sieve, that is, including 10-membered or 12-membered rings, respectively. Examples of suitable molecular sieves include MFI, MEL, MFI / MEL co-biomass, MTW, TUN, UZM-39, IMF, UZM-44, UZM-54, MWW, UZM-37, UZM-8, UZM-8HS. Examples of suitable molecular sieves also include FER, AHT, AEL (SAPO-11), AFO (SAPO-41), MRE, MFS, EUO-1, TON (ZSM-22), MTT (ZSM-23), and UZM-53. Additional molecular sieves with larger pores include FAU, EMT, FAU / EMT co-biomass, UZM-14, MOR, BEA, UZM-50, MTW, ZSM-12. Additional examples include MSE and UZM-35.

[0034] MFI is a suitable NEP catalyst. It should be understood that ZSM-5 is an aluminosilicate zeolite belonging to the MFI type of the pentasil zeolite family and has a chemical formula of Na n A1 n Si 96 -nO 192 16H2O (0 < n < 10). In various embodiments, the ZSM-5 zeolite may have a silica / alumina molar ratio of 20 to 1000, 20 to 800, 20 to 600, 25 to 400, 25 to 200, or 25 to 80. In various embodiments, the ZSM-5 zeolite may have a crystal size in the range of 10 nm to 600 nm, 20 nm to 500 nm, 30 nm to 450 nm, 40 nm to 400 nm, or 50 nm to 300 nm.

[0035] NEP catalysts may include bonded zeolites. Binders may include oxides of aluminum, siloxanes, zinc, titanium, zirconium, and mixtures thereof. In embodiments, the binder may include phosphates in the binder or phosphates of the aforementioned oxide binder materials. Preferably, the binder is aluminum phosphate. MFI zeolites may be supported in an alumina-containing binder (such as aluminum phosphate).

[0036] MFI zeolite slurry can be first mixed with a binder and a gelling agent in the form of a colloidal suspension (sol), and then dropped into hot oil to form controlled spheres, producing calcined supports with diameters of 1 / 64 inch to 1 / 10 inch. The spheres can be washed with ammonia to remove sodium ions from the zeolite, dried, and calcined to remove the organic structure-directing agent (OSDA) from the synthesized zeolite. Optionally, the calcined support can be subjected to ammonium ion exchange using an ammonium nitrate solution to remove residual sodium ions, and then dried at 110°C.

[0037] In one embodiment, the NEP catalyst includes a metal on the catalyst. This metal may include transition metals. In another embodiment, the metal may include platinum, palladium, iridium, rhenium, ruthenium, and mixtures thereof. This metal may be a noble metal. In an additional embodiment, a modifier metal may also be included on the catalyst. The modifier metal may include tin, germanium, gallium, indium, thallium, zinc, silver, and mixtures thereof. The modifier metal should be more concentrated on the binder than on the zeolite. In one embodiment, 0.01 wt% to 5 wt% of each of the transition metal and the modifier metal is present on the catalyst. The catalyst may include 0.1 wt% to 3 wt% of the transition metal.

[0038] Metals can be incorporated into the binder through evaporation impregnation. A platinum solution (such as tetraamine platinum nitrate or chloroplatinic acid) can be contacted with a spherical support that has been calcined and ion-exchanged in a rotary evaporator, followed by drying and oxidation.

[0039] In a preferred embodiment, the NEP catalyst comprises metal on a spherical support for the catalyst. Preferably, the metal on the binder is greater than the metal on the zeolite. In an embodiment, at least 60% by weight, suitably at least 70% by weight, preferably at least 80% by weight, and most preferably at least 90% by weight of metal are on the binder. In an embodiment, the zeolite and / or the entire NEP catalyst is steam-oxidized to remove the metal from the zeolite. Steaming is preferably carried out after the metal has been added to the catalyst. The dried, metal-impregnated binder support can be steam-oxidized in air for a sufficient time to provide the NEP catalyst. Steam oxidation in air at a temperature of 500°C to 650°C and 5 mol% to 30 mol% of steam for 1 to 3 hours may be suitable.

[0040] NEP catalysts are preferably reduced to activate them for catalyzing the NEP reaction. For example, the catalyst may be reduced in a stream of hydrogen at 500°C to 550°C for 2 to 5 hours prior to contact with the feed.

[0041] Following alkane conversion, a light alkane stream is discharged from NEP reactor 16 into effluent line 18. The light alkane stream may contain at least 40% by weight ethane or at least 40% by weight propane, or at least 70% by weight, and preferably at least 80% by weight, ethane and propane. The ethane to propane ratio may range from 0.1 to 5. The light alkane stream may contain less than 15% by weight, suitably less than 12% by weight, more suitably less than 10% by weight, preferably less than 8% by weight, more preferably less than 6% by weight, and most preferably less than 5% by weight of methane.

[0042] It has been found that the presence of sulfur in the NEP reaction does not significantly affect the conversion rate. At least 5 W ppm of sulfur can be present in the feed without significantly affecting the conversion rate. The NEP catalyst is expected to be able to handle up to 100 W ppm, and possibly 200 W ppm, of sulfur without significantly affecting the conversion rate.

[0043] The light alkane stream can be cooled and fed to NEP separation unit 20. NEP separation unit 20 can be a fractionating column or a series of fractionating columns and other separation units that separate the light alkane stream in line 18 into a hydrogen stream in line 22, an ethane stream in line 24, a propane stream in line 26, and a heavy stream in line 12. NEP separation unit 20 may include a demethanizer that separates the light alkane stream into a gas stream in the top line and a C2+ alkane stream in the bottom line. The gas stream can be sent to a hydrogen purification unit, such as a PSA unit, to recover the hydrogen in line 22 for use in the NEP reactor 16. Residual methane from the hydrogen purification unit can be used as fuel gas. The C2+ alkane stream can then be fed to a deethanizer to produce an ethane stream in the top line 24 and a C3+ alkane stream in the bottom line. The C3+ alkane stream can then be fed into a propane stripper to generate a propane stream in the top line 26 of the propane stripper and a heavy alkane stream, which may include C4+ hydrocarbons, in the recovery line 12. The NEP separation unit 20 may be in other forms.

[0044] For example, the NEP separation unit 20 may omit the demethanizer, and the light alkane stream in line 18 may be fed to a deethanizer, which generates a C2 stream in the top line of the deethanizer. The C2 stream can be separated in a hydrogen purification unit to recover the hydrogen stream in line 22, while residual ethane and methane from the hydrogen purification unit may include or supplement the ethane stream in line 24. The hydrogen purification unit may include a membrane unit, and the hydrogen recovered from the membrane unit may be further purified in an absorber and then recovered in line 22 to the NEP reactor 16. In an additional alternative, the C2 stream from the deethanizer may be fed into an ethylene production unit 30, where ethane is converted to ethylene, but methane and hydrogen are inertly passed through for recovery in a downstream ethylene recovery unit.

[0045] The ethane stream in line 24 can be fed into ethylene production unit 30, where the ethane in the ethane stream is converted into ethylene. In an embodiment, ethylene production unit 30 is a steam cracking unit. The ethane stream in line 24 can be cracked in a furnace under steam to produce a cracked stream including ethylene stream 32. The ethane stream can be fed into the ethane steam cracking unit in the gas phase. The ethane steam cracking unit can preferably be operated at a temperature of 750°C (1382°F) to 950°C (1742°F). The cracked stream leaving the furnace of the ethane steam cracking unit can be in a superheated state. One or more quench towers or other devices known in the art, but preferably oil quench towers and / or water quench towers, can be used to quench the cracked stream or separate it into multiple cracked streams. The ethane steam cracking unit may also include additional distillation towers, amine scrubbing towers, compressors, expanders, etc., to separate the cracked stream into cracked streams rich in individual light olefins, wherein the main component is the ethylene stream in line 32. Based on the ethane stream in line 24, the ethylene stream may include an ethylene yield of at least 75% by weight, preferably at least 80% by weight. Among the other components in the cracked stream exiting the ethane steam cracker, the ethylene production unit 30 may contain hydrogen, methane, propylene, butene, and pyrolysis gases. Each of these components may be recycled and further processed.

[0046] The ethylene stream in line 32 and the propylene stream from ethylene production unit 30 can be recycled or transported to a polymerization unit, chemical plant, or exported. The butene stream can be recycled and used to produce plastics or other petrochemical products by methods such as polymerization, or exported. Based on the ethane stream in line 32, product recovery of at least 50% by weight, typically at least 60% by weight, and suitably at least 70% by weight of valuable ethylene, propylene, and butene products can be achieved from ethane steam cracking unit 30.

[0047] In another embodiment, ethylene production unit 30 may be an oxidative dehydrogenation (ODH) unit. An ethane stream from line 24 may be fed into the ODH unit. The ethane ODH method is an alternative technique for the conversion of ethane to ethylene via ethane steam cracking or ethane pyrolysis. Ethane ODH involves contacting an ethane feed and an oxygen source in the presence of an ODH catalyst in the ODH reaction zone under conditions of at least a portion of the oxidative dehydrogenation of ethane to produce a product stream comprising ethylene, carbon oxides, water and unreacted oxygen, acetic acid and other organic acids, and unconverted ethane. The oxygen source may be an oxygen-containing stream or an oxygen-containing material, such as a metal oxide. Mixed metal oxide catalysts have been found to be effective for oxidative dehydrogenation reactions.

[0048] The ODH reactor can use a mixed metal oxide catalyst and operates at temperatures of 300°C to 400°C, producing over 90% ethylene and acetic acid, both of which are useful products. The ODH unit produces an ethylene stream and an acetic acid stream in line 32. Ethane oxidative dehydrogenation using a mixed metal oxide catalyst can be performed at temperatures of 300°C to 500°C, preferably 350°C to 450°C, at pressures of 0.1 bar to 20 bar, preferably 0.1 bar to 10 bar, and space velocities of 1000 cm³ / (gcat·hr) to 5000 cm³ / (gcat·hr), wherein the molar ratio of ethane to oxygen is 1.5:1 to 2:1, with sufficient inert diluent to achieve safe operating conditions. MoVNbTe oxide catalysts (and other related materials) with an M1-type structure are believed to be the optimal catalysts for ethane ODH.

[0049] Ethane ODH using a mixed metal oxide catalyst (where the catalyst is also an oxygen source) can be performed at temperatures of 600°C to 900°C, preferably 750°C to 850°C, at pressures of 0.1 bar to 20 bar, preferably 0.1 bar to 10 bar, and gas hourly space velocities of 1000 h⁻¹ to 5000 h⁻¹. Mg₆MnO₈ oxide catalysts (and other related materials) have been identified as preferred materials for the ethane ODH process, where the catalyst is also an oxygen source.

[0050] The propane stream in line 26 can be fed into propylene production unit 40, where the propane in the propane stream is converted into propylene. Propylene production unit 40 can be an alkane dehydrogenation (PDH) unit. PDH catalysts are used in dehydrogenation reaction methods to catalyze the dehydrogenation of alkanes such as propane. Conditions in the dehydrogenation reactor may include a temperature of 500°C to 800°C, a pressure of 40 kPa to 310 kPa, and a catalyst-to-oil ratio of 5 to 100.

[0051] The dehydrogenation reaction can be carried out in a fluidized manner, where a gas containing reactant alkanes, with or without a fluidized inert gas, is distributed to the reactor in a way that elevates the dehydrogenation catalyst while simultaneously catalyzing the dehydrogenation of the alkanes. During the catalytic dehydrogenation reaction, coke deposits on the dehydrogenation catalyst, resulting in a decrease in catalyst activity. The dehydrogenation catalyst must then be regenerated in a regenerator. The regenerator combusts the coke from the dehydrogenation catalyst and fuel gas to ensure sufficient enthalpy in the dehydrogenation reactor to promote the endothermic reaction.

[0052] The selected dehydrogenation catalyst should minimize cracking and favor dehydrogenation. Catalysts suitable for this study include active metals that can be dispersed in porous inorganic support materials such as silica, alumina, aluminosilicate, zirconium oxide, or clay. Exemplary embodiments of the catalyst include alumina or silica-alumina containing gallium, noble metals, and alkali or alkaline earth metals.

[0053] The catalyst support comprises a support material, a binder, and optional filler materials to provide physical strength and integrity. The support material may include alumina or silica-alumina. Silica sol or alumina sol may be used as a binder. The alumina or silica-alumina typically comprises alumina with γ, θ, and / or δ phases. The nominal diameter of the catalyst support particles can be from 400 micrometers to 5000 micrometers, with an average diameter of 600 micrometers to 3500 micrometers. Preferably, the surface area of ​​the catalyst support is from 85 m² / g to 140 m² / g.

[0054] Fluidized dehydrogenation catalysts may include a dehydrogenating metal on a support. The dehydrogenating metal may be one or a combination of transition metals. Noble metals may be preferred dehydrogenating metals, such as platinum or palladium. Gallium is an effective metal for the dehydrogenation of alkanes. The metal may be deposited on the catalyst support by impregnation or other suitable methods, or included in the support material or binder during catalyst preparation.

[0055] The acid functionality of the catalyst should be minimized to prevent cracking and facilitate dehydrogenation. Alkali metals and alkaline earth metals can also be included in the catalyst to reduce its acidity. Rare earth metals can be included in the catalyst to control its activity. Metals can be incorporated into the dehydrogenation catalyst at concentrations from 0.001 wt% to 10 wt%. In the case of noble metals, it is preferable to use noble metals at concentrations from 10 parts per million (ppm) to 600 ppm by weight. More preferably, it is preferable to use noble metals at concentrations from 10 ppm to 100 ppm by weight. Platinum is a preferred noble metal. Gallium should be present in the range of 0.3 wt% to 3 wt%, preferably 0.5 wt% to 2 wt%. Alkali metals and alkaline earth metals can be present in the range of 0.05 wt% to 1 wt%.

[0056] The regenerated catalyst can be contacted with the propane stream in line 26, and perhaps with a fluidizing gas, to elevate the propane stream and dehydrogenation catalyst into the riser during dehydrogenation. Above the riser, spent dehydrogenation catalyst and propylene products can be separated by a centrifugal separator. The propylene product gas can be quenched with a cooling fluid to prevent over-reaction and the formation of undesirable byproducts. Separation of the propylene products may include quenching contact and fractionation to produce a propylene product stream in line 42. Unreacted propane can be recovered to the dehydrogenation reactor, and the light gas can be recovered as fuel gas to the regenerator for combustion, thereby providing enthalpy for the reaction.

[0057] Propylene production units may also employ catalytic moving bed reactors. The reactor section may comprise several radially flowing reactors in parallel or series, heated by a charge and interstage heaters. Propane, possibly with added hydrogen, flows from a screened central tube through an annular dehydrogenation catalyst bed to an outer effluent ring in each dehydrogenation reactor. The flow may be reversed. The dehydrogenation catalyst may comprise noble metals and mixtures thereof, modifiers selected from the group consisting of alkali metals or alkaline earth metals and combinations thereof, components selected from the group consisting of tin, germanium, lead, indium, gallium, thallium and combinations thereof, and porous supports forming catalyst particles. The catalyst support may comprise oil-droplet alumina spheres.

[0058] Dehydrogenation conditions may include temperatures ranging from 400°C to 900°C, pressures from 0.01 atm to 10 atm, and liquid hourly space velocities (LHSVs) from 0.1 hr⁻¹ to 100 hr⁻¹. The pressure in the dehydrogenation reactor should be maintained at the lowest feasible level, consistent with equipment limitations, to maximize the chemical equilibrium advantage. Waste dehydrogenation catalyst in the annular catalyst bed can be removed from the bottom of the bed and transferred to a regenerator for combustion of coke from the catalyst in air at 450°C to 600°C. Noble metals on the catalyst can be redispersed, dried, and returned to the top of the dehydrogenation catalyst bed as regenerated dehydrogenation catalyst via oxyhalogenation.

[0059] The dehydrogenated effluent from propylene production unit 40 is cooled, compressed, and dried, and hydrogen is cryogenically separated from the hydrocarbons, with a net gas purity of 85 mol% to 93 mol%. The hydrocarbon liquid is selectively hydrogenated to convert dienes and acetylene, and the hydrocarbon liquid is fractionated in a deethaner column to remove ethane, and propylene is separated from propane in a propane-propylene separator to provide polymer-grade propylene in line 42. Propane can be recovered as feed to propylene production unit 40.

[0060] The heavy stream in line 12, which can be removed from the bottom of the propane stripper, may include C4+ alkanes. The heavy stream in line 12 can be recycled to NEP reactor 16 by combining it with the alkane stream in line 10 and the hydrogen stream in line 22, and then loaded into NEP reactor in line 11 to produce more ethane and propane.

[0061] In an alternative embodiment, the light alkane stream in line 18 may be separated by NEP separation unit 20 into a hydrogen stream in line 22, an ethane stream in line 24, and a heavy stream in line 12, as previously described, but the propane stream in line 26 may include isobutane. In this embodiment, the propane and isobutane streams in line 26 may be fed to propylene production unit 40. Propylene production unit 40 may be equipped to dehydrogenate propane in the propane and isobutane streams 26 to propylene in the same dehydrogenation reactor, and to dehydrogenate isobutane in the propane and isobutane streams 26 to isobutene. The fractionation section from propylene production unit 40 may include a depropanizer downstream of the bottom line of the propylene-propane separator to separate unreacted propane from C4 hydrocarbons at the top of the separator. The isobutylene-isobutane separation tower, which is connected downstream of the bottom of the propane depropanizer, can provide isobutylene that can be recovered as a product at the top of the tower, and the unreacted isobutane at the bottom of the separator can be returned to the propylene production unit 40 to be converted into isobutylene.

[0062] Unreacted C4+ hydrocarbons, including C4 and higher alkanes, in the heavy stream 12 can be recovered to be combined with the naphtha stream in feed line 10 and the hydrogen stream in line 22, and fed into the NEP reactor 16 in line 11.

[0063] Figure 2 It shows Figure 1 An alternative embodiment of the implementation scheme employs the n-butene conversion unit 50. It has the same characteristics as... Figure 1 The same configuration Figure 2 The components in will have the same Figure 1 The same icon number as in the text. It has the same... Figure 1 Different configurations of corresponding elements Figure 2 The elements in the diagram have the same reference numerals, but are indicated by an apostrophe ('). Figure 2 The configuration and operation of the implementation scheme are basically the same as Figure 1 The same as in the above, but with the following exceptions.

[0064] In addition to the hydrogen stream in line 22, the ethane stream in line 24, and the propane and isobutane streams in line 26, NEP separation unit 20' also provides a n-butane product stream in line 46. The n-butane stream in line 46 can be fed into butene production unit 50 to convert n-butane in the n-butane stream into n-butene. This embodiment may be useful where the propylene production unit cannot process propane and n-butane in the same dehydrogenation reactor. Butene production unit 50 may include, as for... Figure 1 The dehydrogenation unit described in the propylene production unit of the implementation scheme.

[0065] The butene production unit 50 can feed the n-butene stream to the n-butene-n-butane separation tower. The n-butene stream in the top pipeline of the separator tower can be regarded as the n-butene product in pipeline 52, while the unreacted n-butane in the bottom pipeline can be recovered and returned to the dehydrogenation unit in the butene production unit 50 for conversion into n-butene.

[0066] In some cases, butene production unit 50 may be a butene production unit capable of converting both n-butane and isobutane into n-butene and isobutene. In this case, line 46 carries a propane-lean butane stream to butene production unit 50, and the propane stream in line 26 carries a butane-lean propane stream to propylene production unit. The butene production unit will produce a butene product stream in line 52 that may include isobutene and butene.

[0067] The heavy stream from NEP separation unit 20' in line 12 may include a C5+ hydrocarbon stream, but it may include C4 hydrocarbons and also include a C4+ hydrocarbon stream. Unreacted C4+ or C5+ hydrocarbons in the heavy stream 12, including C4 or C5 and heavier hydrocarbons, may be recovered and combined with the naphtha stream in feed line 10 and the hydrogen stream in line 22, and fed into the NEP reactor 16 in line 11.

[0068] It was also found that the heavy alkane stream in pipeline 12 may also include aromatics, benzene, toluene, and xylene. Figure 3 It shows Figure 1 An alternative embodiment of the implementation scheme, which maximizes the yield of ethane and propane by hydrotreating the entire heavy stream in line 12'' of the hydrotreating reactor 60 to saturate the aromatic rings to cycloalkanes, and then supplying a recovery feed to the NEP reactor 16 in line 62. Figure 1 The same configuration Figure 3 The components in will have the same Figure 1 The same icon number as in the text. It has the same... Figure 1 Different configurations of corresponding elements Figure 3 The elements in the diagram will have the same reference numerals, but will be indicated by double apostrophes (''). Figure 3 The configuration and operation of the implementation scheme are basically the same as Figure 1 The same as in the above, but with the following exceptions.

[0069] The heavy stream, including C4+ alkanes and aromatics, benzene, toluene and xylene, in pipeline 12" is mixed with the hydrogen stream in pipeline 64, heated and loaded into the hydrotreating reactor 60.

[0070] The hydrotreatment reactor 60 may have one or more hydrotreatment catalyst beds to saturate the aromatic rings in the heavy stream. The heavy stream can be fed into the hydrotreatment reactor 60, and the hydrotreatment inlet temperature can be in the range of 200°C (392℉) to 400°C (752℉). If more than one catalyst bed is used, the hydrotreatment reactor 60 may employ an interlayer hydrogen quenching flow.

[0071] Suitable hydrotreating catalysts are any known conventional hydrotreating catalysts, and include those composed of at least one Group VIII metal (preferably iron, cobalt, and nickel, more preferably cobalt and / or nickel) and at least one Group VI metal (preferably molybdenum and tungsten) on a high surface area support material (preferably alumina). Other suitable hydrotreating catalysts include zeolite catalysts, and noble metal catalysts, wherein the noble metal is selected from palladium and platinum. Within the scope of this specification, more than one type of hydrotreating catalyst may be used in the same hydrotreating reactor 60. Group VIII metals are typically present in amounts ranging from 2% to 20% by weight, preferably from 4% to 12% by weight. Group VI metals will typically be present in amounts ranging from 1% to 25% by weight, preferably from 2% to 25% by weight. Typically, hydrotreating conditions include pressures ranging from 700 kPa (100 psig) to 21 MPa (3000 psig). The hydrotreating outlet temperature may be in the range of 300°C (572°F) to 427°C (800°F).

[0072] Saturated aromatics and C4+ alkanes in the hydrotreated heavy stream of line 62 can be recovered and added to the naphtha stream of line 10 and the hydrogen stream of line 22, and then loaded into NEP reactor 16 in line 11. Hydrogen from hydrotreatment reactor 60 can be recovered along with saturated aromatics to NEP reactor 16 to reduce or eliminate the hydrogen demand from line 22.

[0073] In another embodiment, the NEP separation unit 20+ includes a butanizer that separates C4 and possibly C5 hydrocarbons for recovery into the NEP reactor 16 in the pipeline 12+, while retaining aromatics for further processing and value determination. Figure 4 It shows Figure 1 An alternative implementation scheme that retains aromatics. Having the same... Figure 1 The same configuration Figure 4 The components in will have the same Figure 1 The same icon number as in the text. It has the same... Figure 1 Different configurations of corresponding elements Figure 4 The components in the diagram will have the same reference numerals, but will be indicated by a plus sign (+). Figure 4 The operation and structure of the implementation plan are basically the same as Figure 1The same applies, but with the following exceptions.

[0074] NEP separation unit 20+ includes a butanizer that separates a heavy stream containing C4 and possibly C5 alkanes from the propaneizer bottom stream in the butanizer overhead line for recovery to NEP reactor 16 in line 12+ and separates C5+ or C6+ aromatics in the butanizer bottom line 28. The aromatics in the butanizer bottom line 28 can be further processed for valuation of valuable aromatics.

[0075] The aforementioned disclosure provides a method for converting naphtha into ethane and propane feedstocks, wherein the conversion to ethylene and propylene is maximized in ethylene and propylene production units, respectively. The method provided herein produces an increased amount of ethane, thus enabling the production of even higher amounts of ethylene.

[0076] Example 1 and Example 2

[0077] Zeolite MFI with a Si / Al2 ratio of 40 was synthesized by crystallization of a mixture consisting of silica, alumina, alkali metals, and OSDA containing tetrapropylammonium (TPA). The synthesized zeolite MFI was first mixed with AlPO4 sol and a gelling agent, and the resulting slurry was dripped into hot oil. The size of the spheres was controlled to obtain calcined supports with a diameter of 1 / 16 inch. The spheres exiting the pelletizing section were separated from the oil and conveyed to a washing section for ammonia washing to remove sodium ions from the zeolite. The wet spheres were dried and then calcined to remove OSDA. Optionally, the calcined supports were subjected to NH4 ion exchange with an ammonium nitrate solution to remove residual sodium ions, followed by a drying step at 110°C prior to the metal incorporation step. Platinum incorporation was performed by evaporative impregnation, via contacting the supports with a solution of a platinum-containing compound (such as tetraamine platinum nitrate) in a rotary evaporator, followed by drying with heat introduced from a steam jacket. Once the spheres are freely rolled, the impregnated support is steam oxidized in a quartz tube in a Zone 3 furnace at 607°C for 2 hours in flowing air with 7.7 mol% or 14.3 mol% steam. The target platinum levels on the finished catalyst are 0.005 wt%, 0.02 wt%, 0.10 wt%, and 0.40 wt%, and are designated as Comparative Examples and Examples 1.1 to 1.4 (for 7.7% steam) and 2.1 to 2.4 (for 14.3% steam), as shown in Table 2.

[0078] Table 2

[0079]

[0080] Example 3

[0081] A series of 1 / 8-inch trefoil extrudates were prepared using a formulation of 65% MFI-40 prepared according to Example 1, and a total of 35% Hisil 532 SiO2 and Ludox SiO2. 1% Methocel A extrusion aid was incorporated into the formulation. The extrudates were dried at 110°C and calcined at 520°C in flowing air in a muffle furnace. The calcined support was subjected to ammonium ion exchange using a 1.0M ammonium nitrate solution to remove residual sodium. The ammonium ion-exchanged support was dried at 110°C, and then chloroplatinic acid (CPA) was incorporated into the support by evaporative impregnation to obtain 0.2% platinum on the finished catalyst. The impregnated support was steam-oxidized for 2 hours at a temperature range of 565°C to 650°C and a vapor concentration range of 7.7% to 20%. As shown in Table 3, the finished catalysts were designated as Comparative Examples and Examples 3.1 to 3.11. Available platinum (or platinum dispersion) was determined by hydrogen chemisorption via a volumetric vacuum method substantially according to ASTM D3908-20. Available chemisorption sites were reported as μmol H2 per gram of catalyst, and platinum dispersion was calculated using the amount of chemisorbed H2 and the platinum content of the catalyst. The acidity of the catalyst was determined by ammonia-programmed temperature desorption (NH3-TPD), or calculated based on the correlation between NH3-TPDs and the zeolite framework composition measured by infrared (stretch frequency) technology. Acidity measurements via NH3-TPD were performed using a Micromeritics AutoChem 2920 equipped with a TCD. In this acidity measurement, 250 mg–600 mg was loaded, and pretreatment was performed by increasing the temperature from 25 °C to 650 °C at 10 °C / min in a 100 mL stream of 20 / 80 O2 / helium and holding at 650 °C for 1 hour. Once cooled to 150°C, NH3 adsorption was performed by exposing the sample to NH3 flowing at 150°C. This was accomplished by purging several quantitative loops (volume of each loop) of a 10 / 90 NH3 / helium mixture into the sample chamber using a helium flow of 20 ml / min. Subsequently, a helium flow of 75 ml / min was passed through the sample chamber to purge NH3 that was not adsorbed or loosely adsorbed onto the sample at 150°C until the effluent composition remained unchanged. After NH3 purging, NH3 desorption was performed by increasing the sample chamber temperature from 150°C to 650°C using a helium flow of 75 ml / min at 10°C / min. The amount of NH3 removed from the sample was monitored using a TCD, and the results were reported as micromoles of NH3 desorbed per gram within the specified temperature range. The acidity obtained by NH3-TPD is characterized by a low-temperature peak at approximately 250°C and a high-temperature peak at approximately 400°C, and the acidity associated with the high-temperature peak is reported in the table. When NH3-TPD results are missing, NH3-TPD acidity is obtained based on its correlation with the skeleton composition (acidity) obtained by infrared spectroscopy as described below.

[0082] The skeletal composition (SiO2 and Al2O3) obtained by infrared spectroscopy was determined according to the procedure described herein. First, the sample was ground into a fine powder using an agate mortar and pestle. The powder was mixed with XL ultrapure spectral grade KBr powder to obtain a dilution of 1000:1. Granules with a diameter of 13 mm were prepared by pressing 250 mg of the sample / KBr mixture at 10,000 lbs using a hydraulic press. The KBr granules were dried in a drying oven at 160 °C for >2 hours. The skeletal composition (SiO2 and Al2O3) was determined using a Nicolet is-50 FTIR spectrometer with a DTGS detector at a 2 cm⁻¹ aperture. -1 Spectra were collected in transmission mode with 128 scans performed. Standard β-zeolite was used as the daily reference material. The positions of the asymmetric TOT (T=Si,Al) bands in zeolites were negatively correlated with the framework Al content in many zeolites.

[0083] Table 3

[0084]

[0085] Example 4

[0086] 1 / 16-inch cylindrical extrudates were prepared using a formulation of 65% MFI-40 prepared according to Example 1, and a total of 35% Hisil 532 SiO2 and Ludox AS-40 SiO2. 1% Methocel A extrusion aid was incorporated into the formulation. The extrudates were dried at 110°C and calcined at 520°C in flowing air in a muffle furnace. The calcined support was subjected to ammonium ion exchange using a 1.0M ammonium nitrate solution to remove residual sodium. The ammonium ion-exchanged support was dried at 110°C, and then tetraamine nitrate platinum or chloroplatinic acid was incorporated into the support by evaporative impregnation to obtain 0.2% or 0.4% platinum on the finished catalyst. The impregnated support was vapor-oxidized at 607°C and 14.3% vapor for 2 hours, and the finished catalysts were named Examples 4.1, 4.2, and 4.3, respectively.

[0087] Preparation of Comparative Example 4.4. Using an alternative preparation sequence as in Example 2.4, the oil droplet spheres incorporating AlPO4 were first calcined, subjected to ammonium ion exchange to remove residual sodium, and then steam-calcined at 607°C for 2 hours using 14.3 mol% steam. The steam-calcined support was impregnated with a tetraamine nitrate platinumate solution by evaporation, followed by dry oxidation at 520°C for 2 hours in flowing air. This catalyst contained 0.4 wt% platinum and was named Comparative Example 4.4. Comparative Example 4.5 was prepared according to the same procedure as in Example 4.2, except that steam oxidation was replaced by dry oxidation at 520°C. Examples 4.6 and 4.7 were prepared according to the procedures described for Examples 4.3 and 4.2, except that MFI-40 was replaced with MFI-23 obtained from Zeolyst. Comparative Example 4.8 was prepared by first preparing an MFI-23 support incorporating silica using oil droplet spheres. The oil droplet carrier was post-treated by calcining a dry substrate, ion-exchanging it with ammonium nitrate solution, and drying it in a muffle furnace at 110°C. The ammonium ion-exchanged carrier was impregnated with tetraamine platinum(II) nitrate to achieve a target of 0.2% platinum, followed by steam oxidation at 607°C using flowing air containing 14.4% H2O. The carrier used in Example 4.9 was prepared according to the procedure described for the carriers of Examples 4.1-4.3, but with a trefoil shape and a 1 / 8-inch circumference. Platinum doping and subsequent post-treatment followed the same procedure as those for Example 4.1, using steam oxidation conditions of 7.7% steam at 565°C for 2 hours. Example 4.10 was obtained by first preparing alumina phosphate particles containing platinum (0.36 wt%) by impregnating them with chloroplatinic acid solution, then combining them with MFI-40 particles, followed by granulation (to 20 × 60 mesh) and steam oxidation at 565°C with 7.7% steam for 2 hours. Example 4.11 was obtained as follows: First, platinum-containing particles (0.36 wt%) were prepared by impregnating MFI-40 particles with a chloroplatinic acid solution. These particles were then combined with alumina phosphate particles, granulated (to 20 × 60 mesh), and steam-oxidized at 7.7% steam and 565°C for 2 hours. Example 4.12 was obtained as follows: First, platinum-containing particles (0.36 wt%) were prepared by impregnating SiO2 particles with a chloroplatinic acid solution. These particles were then combined with MFI-40 particles, granulated (to 20 × 60 mesh), and steam-oxidized at 7.7% steam and 565°C for 2 hours.

[0088] The descriptions and characteristics of the catalyst examples and comparative examples are summarized in Table 4.

[0089] Table 4

[0090]

[0091] Example 5

[0092] Catalytic testing was conducted in a piston flow tubular reactor using a setup consisting of multiple parallel reactors. The composition of the feed blend, shown in Table 5, simulates the composition of light naphtha. The finished catalyst was crushed, and 1 ml of 20×60 mesh catalyst was charged into the reactor. The catalyst was pre-reduced in a flowing hydrogen stream at 520°C for 3 hours, and then lowered to the target reaction temperature before the feed was introduced. The reaction was carried out at a pressure of 6 bar, a hydrogen to hydrocarbon molar ratio of 2.0, and a reaction time of 2.0 h. -1 Catalytic performance was measured at a weight hourly space velocity (WHSV) within a temperature range of 390°C to 500°C. Results at 470°C are shown in Table 6. The results indicate that the production of ethane, propane, and combinations of ethane and propane will require a minimum amount of platinum of 0.02% by weight, preferably 0.1% by weight, and most preferably 0.2% by weight.

[0093] Table 5 Simulated commercial light naphtha

[0094]

[0095] Table 6.1

[0096]

[0097] Table 6.2

[0098]

[0099] Example 6

[0100] The catalyst from Example 3 was tested in a 7 / 8-inch diameter plug flow reactor. Catalysts from 1 / 16-inch or 1 / 8-inch diameter extrusions were premixed with α-alumina and then charged into the reactor. The feed blend consisted of 30% simulated light naphtha as shown in Table 5 and 70% aromatics primarily composed of toluene. The catalyst was pre-reduced in a flowing hydrogen stream at 520°C for 3 hours and then cooled to the target reaction temperature before being introduced into the feed. Catalytic performance was measured at 21 bar pressure and a hydrogen to hydrocarbon (including light naphtha and aromatics) molar ratio of 1.0 over a temperature range of 400°C to 500°C (set temperature). The results at 475°C are shown in Table 7. Clearly, to produce ethane and propane in a combination of greater than 80% and an ethane to propane ratio greater than 0.5, a minimum acidity greater than 0.12 mmol / g, and preferably greater than 0.15 mmol / g, will be required.

[0101] Table 7

[0102]

[0103] Example 7

[0104] The catalytic performance of catalyst Example 4 was measured using the equipment and testing protocol described in Example 5. The test results are reported as the amount of methane formed at a 55% ethane yield, as shown in Table 8. Performance of selective catalysts Examples 1 and 2 is also included. The results show that, at a given ethane yield, methane formation decreases with decreasing hydrogen chemisorption and decreasing ratio of chemisorbed hydrogen to platinum, as shown in Table 8. Depending on the binder used to produce the catalyst, it is desirable that the H2 chemisorption is less than 7 μmol / g, preferably less than 6 μmol / g, and most preferably less than 5 μmol / g. It is desirable that the ratio of chemisorbed hydrogen to platinum is less than 0.5, preferably less than 0.45, and most preferably less than 0.35.

[0105] Table 8

[0106]

[0107] Example 8.1

[0108] The 1 / 8:trilobal extrudate prepared according to Example 3 was used for catalyst refining in Example 8. 50 g of the calcined extrudate was impregnated with a solution of 190 mg tetraammineplatinum(II) nitrate (TAPN) dissolved in 77 g of water. In the metal incorporation, the calcined extrudate was first added to a rotary impregnator, along with a Pt solution. The impregnator was first cold-rolled at room temperature for 1 hour, then hot-rolled via a steam jacket until the extrudate was freely rolled. The impregnated base material was dried overnight at 110°C. The sample was then steam-oxidized in a Zone 3 furnace at 565°C in flowing air containing 8% steam for 2 hours. This catalyst was designed as in Example 8.1.

[0109] Examples 8.2 and 8.3

[0110] Catalysts were prepared according to the procedure described in Example 8.1 in Examples 8.2 and 8.3, except that 2.0 g and 4.0 g of ammonium nitrate were added to the TAPN impregnation solution, respectively.

[0111] Examples 8.4 and 8.5

[0112] Catalysts were prepared according to the procedure described in Example 8.1 in Examples 8.4 and 8.5, except that 2.0 g and 4.0 g of nitric acid were added to the TAPN impregnation solution, respectively.

[0113] Examples 8.6 to 8.10

[0114] Examples 8.6 to 8.10 were prepared according to Examples 8.1 to 8.5, except that the impregnated base material was dry-oxidized in a Zone 3 furnace at 565°C in flowing air without steam.

[0115] Example 9.1

[0116] 1 / 8-inch trefoil extruders were prepared using 35% Versal Al2O3 and 65% MFI with Si / Al2=40 synthesized according to Example 3. The calcined support was purified with chloroplatinic acid (CPA) using a volume-based solution-to-support ratio to obtain a target of 0.2% by weight of platinum, and then vaporized at 565°C for two hours in flowing air containing 8% vapor.

[0117] Example 9.2

[0118] 25 g of the calcined support (65% MFI-Q / 35% Al2O3) described in Example 9.1 was impregnated with a solution of 1.52 g of CPA solution and 3.0 g of HCl acid dissolved in 36 g of water. The calcined support was first added to a quartz impregnating machine equipped with a steam jacket, followed by the impregnation solution. The impregnating machine was first cold-rolled at room temperature for one hour, then hot-rolled until the extrudate was freely rolled. The sample was dried overnight in a box oven at 110°C. The sample was then steam-oxidized in a Zone 3 furnace at 565°C in flowing air containing 8% steam for 2 hours. This catalyst was designed as in Example 9.2.

[0119] Example 9.3

[0120] Example 9.3 was prepared according to Example 9.1, except that TAPN was used instead of CPA.

[0121] Example 9.4

[0122] 30 g of the calcined support (65% MFI-Q / 35% Al2O3) described in Example 9.1 was impregnated with a solution of 113 mg TAPN and 1.2 g ammonium nitrate dissolved in 50 g of water. The calcined support was first added to a quartz impregnating machine equipped with a steam jacket, followed by the impregnation solution. The impregnating machine was first cold-rolled at room temperature for one hour, then hot-rolled until the extrudate was freely rolled. The sample was dried overnight in a box oven at 110°C. The sample was then steam-oxidized in a Zone 3 furnace at 565°C in flowing air containing 8% steam for 2 hours. This catalyst was designed as in Example 9.4.

[0123] Example 9.5

[0124] 30 g of the calcined support (65% MFI-Q / 35% Al2O3) specified in Example 9.1 was impregnated with a solution of 113 mg TAPN nitrate and 2.49 g ammonium nitrate dissolved in 50 g of water. The sample was dried overnight in a box oven at 110 °C. The sample was then steam-oxidized in a Zone 3 furnace at 565 °C in flowing air containing 8% steam for 2 hours. This catalyst was designed as in Example 9.5.

[0125] Example 10

[0126] The platinum concentration distribution across the cross-sections of the extrudates in Examples 8 and 9 was determined using scanning electron microscopy as described below. Catalyst samples were mounted in an acrylic medium and polished downwards to the cross-section. EDS plots of each catalyst cross-section were collected via a SEM-EDS analysis mount. The collected EDS data for each cross-section were processed into Euclidean zones of equal area from the outer edge to the center. Data from each of these zones were output as the concentration distribution for each catalyst. This data was then plotted, with each zone representing a point on the plot and the set of points representing the catalyst cross-section. The catalysts described in Examples 8 and 9 exhibited varying degrees of platinum homogeneity or surface enrichment. Acidity utilization was calculated from the amount of acidity covered by 87.5% platinum from the exterior of the extrudate. These are reported in Table 9 along with the available platinum surface area, measured substantially according to ASTM D3908-20. The amount of platinum located within the zeolite micropores was measured by infrared spectroscopy using the amount of CO (carbon monoxide) interacting with platinum oxide. In determining the platinum located within the zeolite micropores, the dry oxidation and steam oxidation catalysts prepared according to Examples 8 and 9 were pretreated in flowing air at 520°C for 2 hours. Once cooled in helium, CO adsorption was performed by allowing 0.1% CO in helium to flow over the sample for 5 minutes, equilibrating for 15 minutes, and then purging with helium for 2 minutes. The spectra were then acquired. (Source: 2090 cm⁻¹) -1 Up to 2050cm -1 The peak was identified as CO interacting with platinum, which in turn interacts with skeletal oxygen within the zeolite micropores. The amount of platinum located within the zeolite micropores is reported in Table 9.

[0127] Example 10.

[0128] Table 9

[0129]

[0130] Example 11

[0131] Catalytic testing was conducted in a plug flow reactor with a diameter of 7 / 8 inch and a feed blend of 75% isopentane and 25% cyclohexane (by weight) loaded with 10 g of 1 / 8 inch diameter trilobal extruder at a weight hourly space velocity of 2 hr⁻¹, a hydrogen to hydrocarbon molar ratio of 2.0, and a pressure of 300 psig. Performance was measured using an online GC equipped with PONA (boiling point) and aromatic (polar) columns at set temperatures of 420 °C, 460 °C, 490 °C, 510 °C, and 490 °C. Results at the 490 °C set temperature are reported in Table 10.

[0132] Table 10.

[0133]

[0134] It should be noted that the production of light alkanes with high ethane yields is achieved using catalysts with chemisorbed H2 concentrations of less than 6 μmol / g or greater than 7 μmol / g. It should also be noted that ethane, and combinations thereof, are produced using catalysts with H2 / Pt ratios less than 0.6 and greater than 0.7. To maximize the conversion of light alkanes and cycloalkanes into ethane, propane, and their combinations, a platinum area of ​​less than 0.055 CO / mg within the zeolite micropores is desirable, as indicated by infrared spectroscopy at 2080 cm⁻¹. -1 Up to 2050cm -1 The CO interaction within the platinum oxide range was measured. As shown in Table 11, the same observations were also noted for the catalytic performance measured at 460 °C.

[0135] Table 11

[0136]

[0137] Specific implementation plan

[0138] While the following description is presented in conjunction with specific embodiments, it should be understood that the description is intended to be illustrative and not to limit the scope of the foregoing description and the appended claims.

[0139] A first embodiment of the present invention is a catalyst for converting naphtha into ethane and propane, wherein the catalyst contains greater than 0.1% by weight of platinum and has an H2 chemisorption capacity of less than 7 μmol / g. An embodiment of the catalyst of the present invention is a zeolite catalyst, wherein the zeolite in the zeolite catalyst comprises a zeolite structure selected from the group consisting of MFI, FER, MEL, UZM-39, UZM-44, UZM-54, MWW, MFS, AEL, MSE, UZM-35, and MTW. An embodiment of the present invention is one, any, or all of the embodiments described in the preceding embodiments to the first embodiment of this paragraph, wherein the catalyst has an H2 chemisorption capacity of less than 6 μmol / g. An embodiment of the present invention is one, any, or all of the embodiments described in the preceding embodiments to the first embodiment of this paragraph, wherein the catalyst has an H2 chemisorption capacity of less than 5 μmol / g. An embodiment of the present invention is one, any, or all of the embodiments described in the preceding embodiments to the first embodiment of this paragraph, wherein the catalyst has a chemisorption hydrogen to platinum ratio of less than 0.5. One embodiment of the present invention is one, any, or all of the embodiments described in the preceding embodiments to the first embodiment of this paragraph, wherein the catalyst has a chemisorbed hydrogen to platinum ratio of less than 0.45. One embodiment of the present invention is one, any, or all of the embodiments described in the preceding embodiments to the first embodiment of this paragraph, wherein the catalyst has a chemisorbed hydrogen to platinum ratio of less than 0.35. One embodiment of the present invention is one, any, or all of the embodiments described in the preceding embodiments to the first embodiment of this paragraph, wherein the catalyst has an acidity greater than 0.15 mmol / g. One embodiment of the present invention is one, any, or all of the embodiments described in the preceding embodiments to the first embodiment of this paragraph, wherein the catalyst contains 0.2% by weight of platinum.

[0140] A second embodiment of the present invention is a catalyst for converting naphtha into ethane and propane, wherein the catalyst contains at least 0.1% by weight of platinum and has a chemisorbed hydrogen to platinum ratio of less than 0.5. One embodiment of the present invention is one, any, or all of the embodiments described in the preceding to the second embodiments of this paragraph, wherein the catalyst is a zeolite catalyst, and wherein the zeolite in the zeolite catalyst comprises a zeolite structure selected from the group consisting of MFI, FER, MEL, UZM-39, UZM-44, UZM-54, MWW, MFS, AEL, MSE, UZM-35, and MTW. One embodiment of the present invention is one, any, or all of the embodiments described in the preceding to the second embodiments of this paragraph, wherein the catalyst contains 0.2% by weight of platinum. One embodiment of the present invention is one, any, or all of the embodiments described in the preceding to the second embodiments of this paragraph, wherein the catalyst has a chemisorbed hydrogen to platinum ratio of less than 0.35. One embodiment of the invention is one, any, or all of the embodiments described in the preceding to the second embodiments of this paragraph, wherein the catalyst has an acidity greater than 0.12 mmol / g. Another embodiment of the invention is one, any, or all of the embodiments described in the preceding to the second embodiments of this paragraph, wherein the catalyst has an acidity greater than 0.15 mmol / g.

[0141] A third embodiment of the present invention is a method for converting naphtha, the method comprising: contacting a naphtha stream with a zeolite catalyst under conditions of dehydrogenating alkanes to olefins, interconverting the olefins to lighter olefins, and hydrogenating the lighter olefins to produce a light alkane stream comprising more than 80% ethane and propane in an ethane-to-propane ratio greater than 0.5, wherein the catalyst contains more than 0.1% by weight of platinum and has an H2 chemisorption capacity of less than 6 μmol / g. An embodiment of the present invention is one, any one, or all of the embodiments of the preceding embodiments to the third embodiment of this paragraph, wherein the catalyst is a zeolite catalyst, and wherein the zeolite in the zeolite catalyst comprises a zeolite structure selected from the group consisting of MFI, FER, MEL, UZM-39, UZM-44, UZM-54, MWW, MFS, AEL, MSE, UZM-35, and MTW. One embodiment of the present invention is one, any, or all of the embodiments described in the preceding to the third embodiments of this paragraph, wherein the catalyst has a chemisorbed hydrogen to platinum ratio of less than 0.5. One embodiment of the present invention is one, any, or all of the embodiments described in the preceding to the third embodiments of this paragraph, wherein the catalyst has an acidity greater than 0.12 mmol / g. One embodiment of the present invention is one, any, or all of the embodiments described in the preceding to the third embodiments of this paragraph, wherein the catalyst has an acidity greater than 0.15 mmol / g. One embodiment of the present invention is one, any, or all of the embodiments described in the preceding to the third embodiments of this paragraph, wherein the light alkane stream comprises 55% ethane. One embodiment of the present invention is one, any, or all of the embodiments described in the preceding to the third embodiments of this paragraph, wherein the pressure ranges from 0.1 MPa to 3 MPa, and the weight hourly space velocity ranges from 0.3 hr. -1 up to 20 hours -1 And the temperature range is 350℃ to 600℃.

[0142] Although no further detailed description has been provided, it is believed that those skilled in the art can make full use of this disclosure by employing the foregoing description and can readily determine the essential characteristics of this disclosure without departing from the spirit and scope of the invention, and can make various changes and modifications to this disclosure to suit various uses and conditions. Therefore, the foregoing preferred specific embodiments should be understood as illustrative only and not as limiting the remainder of this disclosure in any way, and are intended to cover various modifications and equivalent arrangements included within the scope of the appended claims.

[0143] In the foregoing, all temperatures are expressed in degrees Celsius, and all portions and percentages are by weight unless otherwise specified.

Claims

1. A catalyst for converting naphtha into ethane and propane, wherein the catalyst contains more than 0.1% by weight of platinum and has an H2 chemisorption capacity of less than 6 μmol / g or more than 7 μmol / g or a chemisorption hydrogen to platinum ratio of less than 0.

5.

2. The catalyst according to claim 1, wherein the catalyst is a zeolite catalyst, and wherein the zeolite in the zeolite catalyst comprises a zeolite structure selected from the group consisting of MFI, FER, MEL, UZM-39, UZM-44, UZM-54, MWW, MFS, AEL, MSE, UZM-35 and MTW.

3. The catalyst according to claim 1, wherein the catalyst has a chemisorbed H2 to platinum ratio of less than 0.6 or greater than 0.

7.

4. The catalyst according to claim 1, wherein the catalyst has an H2 chemisorption capacity of less than 5 μmol / g or greater than 7 μmol / g.

5. The catalyst according to claim 1, wherein the catalyst has a chemisorbed hydrogen to platinum ratio of less than 0.5 or greater than 0.

75.

6. The catalyst according to claim 1, wherein the catalyst has an acidity greater than 0.12 mmol / g.

7. The catalyst according to claim 1, wherein the catalyst contains more than 0.1% by weight of platinum.

8. A method for converting naphtha, the method comprising: The naphtha stream is contacted with a zeolite catalyst under conditions in which alkanes are dehydrogenated to olefins, the olefins are interconverted to lighter olefins, and the lighter olefins are hydrogenated to produce a light alkanes stream containing more than 75% ethane and propane with an ethane to propane ratio greater than 0.5, wherein the catalyst contains more than 0.1% by weight of platinum and has an H2 chemisorption capacity of less than 6 μmol / g.

9. The method according to claim 8, wherein the catalyst is a zeolite catalyst, and wherein the zeolite in the zeolite catalyst comprises a zeolite structure selected from the group consisting of MFI, FER, MEL, UZM-39, UZM-44, UZM-54, MWW, MFS, AEL, MSE, UZM-35 and MTW.

10. The method of claim 8, wherein the platinum within the zeolite micropores of the catalyst, as measured by infrared spectroscopy, is less than 0.055 CO area / mg.

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