JP2014522891A - Method and system for reducing the olefin content of a Fischer-Tropsch product stream - Google Patents

Abstract

A method is provided for converting synthesis gas to a liquid hydrocarbon mixture useful as a distillate fuel and / or a lubricating base oil. The synthesis gas is contacted with a synthesis gas conversion catalyst comprising a Fischer-Tropsch synthesis component in the upstream catalyst bed, thereby producing an intermediate hydrocarbon mixture containing C _{21 +} normal paraffins and olefins. The intermediate hydrocarbon mixture is then contacted with a hydroisomerization catalyst and an olefin saturation catalyst, resulting in a product containing no more than about 5% by weight C 21 ₊ normal paraffins and no more than about 25% by weight olefins. . The hydroisomerization and olefin saturation catalyst may be in separate beds or may be mixed downstream a single bed of the synthesis gas conversion catalyst.

Description

FIELD The present invention relates to a process for converting synthesis gas to a liquid hydrocarbon mixture useful as a distillate fuel and / or a lubricating base oil, the process contacting the synthesis gas with a catalyst, thereby producing an olefin component and Producing a liquid containing a paraffin component and saturating the olefin component by contacting the liquid with a hydrogenation catalyst.

BACKGROUND Today, the majority of combustible liquid fuels used in the world are derived from crude oil. However, there are some limitations to using crude oil as a fuel source. For example, crude oil is under limited supply.

Alternative sources for developing flammable liquid fuels are desirable. Abundant resources are natural gas. The conversion of natural gas to flammable liquid fuel involves a first step of converting natural gas, mostly methane, to synthesis gas or synthesis gas (syngas), which is a mixture of hydrogen and carbon monoxide. Typically includes. Fischer-Tropsch synthesis is a known means for converting synthesis gas to higher molecular weight hydrocarbon products. Fischer-Tropsch diesel has a very high cetane number, and is effective in reducing the NO _x and particulates from a diesel engine in blends with conventional diesel them to meet more stringent emission standards It becomes possible to do.

Fischer-Tropsch synthesis is often performed under conditions that produce large amounts of C ₂₁₊ wax, also referred to herein as “Fischer-Tropsch wax,” which must be hydrotreated to provide distillate fuel. Often, the wax is hydrocracked to reduce chain length and then hydrotreated to reduce oxygenates and olefins to paraffin. Hydrocracking tends to reduce the chain length of all hydrocarbons in the feed. When the feed contains hydrocarbons that are already in the desired range, eg, distillate fuel range, hydrocracking of these hydrocarbons is not desired.

  As disclosed in co-pending US patent application Ser. No. 12 / 343,534, which is incorporated by reference in its entirety, a hybrid Fischer-Tropsch catalyst, also referred to herein as a hybrid syngas conversion catalyst, is described and the syngas is Can be converted to a hydrocarbon mixture free of solid wax. One advantage of the catalyst-based process is that the absence of a solid wax phase eliminates the need for waxy product separation and hydroprocessing and / or hydrocracking in a separate reactor. It is. Thus, the hydrocarbon product resulting from this improved process can theoretically be blended with crude oil.

  In practice, however, Fischer-Tropsch synthesis produces a high percentage of olefinic hydrocarbons. An olefinic hydrocarbon is defined as a hydrocarbon having one or more double bonds in the molecule. Olefinic or unsaturated hydrocarbons have devastating potential for purification processes, causing problems including intact heaters and preheat train fouling, storage instability and gum deposits. Moreover, apart from diene saturation, olefin hydrogenation is not carried out in crude oil refining. For this reason, synthetic hydrocarbon mixtures must be treated so as to substantially remove unsaturated hydrocarbons prior to blending into crude oil.

  It would be desirable to have a means to convert the synthesis gas to a hydrocarbon mixture free of solid wax with a low percentage of olefins.

Overview According to one aspect, the present invention provides:
A method for converting a synthesis gas to a hydrocarbon mixture comprising sequentially contacting a feed comprising a mixture of hydrogen and carbon monoxide with the following synthesis gas conversion catalyst, hydroisomerization catalyst, and olefin saturation catalyst: There:
A synthesis gas conversion catalyst in the upstream bed, wherein a first intermediate hydrocarbon mixture containing C _{21 +} normal paraffin and olefin is formed on the synthesis gas conversion catalyst;
A hydroisomerization catalyst containing an acidic component and a metal promoter downstream of the intermediate catalyst bed of the upstream catalyst bed, wherein the C _{21 +} normal paraffin of the first intermediate hydrocarbon mixture is a hydroisomerization catalyst. Forming a second intermediate hydrocarbon mixture containing no more than about 5% by weight of _{C21 +} normal paraffins and olefins, and downstream of the intermediate catalyst bed downstream of the intermediate catalyst bed. An olefin saturation catalyst, wherein the olefin is saturated on the olefin saturation catalyst,
Thereby resulting in a final hydrocarbon mixture containing no more than about 5% by weight C 21 ₊ normal paraffins and no more than about 25% by weight olefins.
Relates to said method.

According to another aspect, the present invention provides:
Contacting the feed comprising the mixture of hydrogen and carbon monoxide with the following synthesis gas conversion catalyst, and a mixture of hydroisomerization catalyst and olefin saturation catalyst, in turn, to the hydrocarbon mixture. The method of conversion is:
A synthesis gas conversion catalyst in the upstream bed, wherein a first intermediate hydrocarbon mixture containing C _{21 +} normal paraffin and olefin is formed on the synthesis gas conversion catalyst;
A mixture of a hydroisomerization catalyst and an olefin saturation catalyst containing an acidic component and a metal promoter in the downstream catalyst bed, whereby the C _{21 +} normal paraffin of the intermediate hydrocarbon mixture is hydroisomerized; And the olefin is saturated, thus forming a hydrocarbon mixture containing no more than about 5% by weight C 21 ₊ normal paraffins and no more than about 25% by weight olefins.
Relates to said method.

FIG. 1 is a schematic diagram illustrating a method for converting synthesis gas to liquid hydrocarbons according to an embodiment of the present invention. FIG. 2 is a schematic diagram illustrating a method for converting synthesis gas to liquid hydrocarbons according to another aspect of the present invention. FIG. 3 is a schematic diagram illustrating a method for converting synthesis gas to liquid hydrocarbons according to yet another aspect of the present invention.

DETAILED DESCRIPTION With reference to FIG. 1, a method of synthesizing a liquid hydrocarbon product 30 in the distillate fuel and / or lube base oil range from a feed of synthesis gas 2 is disclosed. Inside the fixed bed reactor, multiple small diameter tubes (not shown) are placed in a cooling medium such as steam or water. In the process, a method is provided for synthesizing a mixture of olefinic and paraffinic hydrocarbons by contacting the synthesis gas in a reactor with a synthesis gas conversion catalyst in a first upstream catalyst bed 4. The hydrocarbon mixture so formed can range from methane to light waxes containing only trace amounts of carbon above 30 (<0.5 wt.%) And linear, Branched and cyclic compounds may be included. As defined herein, the terms “wax” and “solid wax” refer to C _{21 +} normal paraffin. The terms “Fischer-Tropsch wax” and “C _{21 +} wax” are also used here to denote C _{21 +} normal paraffins _{relative to} each other. The hydrocarbon mixture formed in the catalyst bed 4 is then contacted with the second intermediate catalyst bed 6 in the same reactor. The intermediate bed can include a hydrogenation catalyst for hydrogenating olefins and a catalyst for hydroisomerizing linear hydrocarbons. The upstream bed performs synthesis gas conversion while the intermediate bed performs hydroisomerization and optionally hydrocracking. Syngas conversion and subsequent hydroisomerization is expedient in a single reactor under essentially common reaction conditions without the need to supply separate reactors for hydroisomerization and optional hydrocracking. Can be executed well. “Essentially common reaction conditions” means that the temperature of the cooling medium in the reactor 10 is constant within a few degrees Celsius (such as 0 to 3 ° C.) from one point to the other. Means that the pressure can be balanced between the two beds. In some cases, although not preferred, more than one cooling system may be used that utilizes more than one cooling medium physically separated from each other, where the cooling medium may be at different temperatures. . The temperature and pressure of the upstream bed 4 and the intermediate bed 6 can be somewhat different, but advantageously the temperature and pressure of the two beds need not be controlled separately. The bed temperature will depend on the relative exotherm of the reaction progress within them. The exotherm generated by syngas conversion is greater than that generated by hydrocracking; therefore, for a given reactor tube diameter, the average upstream bed temperature will generally be higher than the average intermediate bed temperature. The temperature difference between the beds will depend on various reactor design factors including, but not limited to, the type and temperature of the cooling medium, the diameter of the tubes in the reactor, the rate of gas flow through the reactor, etc. Let's go. For sufficient thermal control, the temperature of the two beds is preferably maintained within about 10 ° C. of the coolant temperature, and therefore the temperature difference between the upstream and intermediate beds is preferably about It is less than 20 ° C and even less than about 10 ° C. The pressure at the end of the upstream bed is equivalent to the pressure at the top of the intermediate bed. This is because the two floors are open to each other. Note that there will be a pressure drop from the top of the upstream bed to the bottom of the intermediate bed. This is because the gas is forced through a narrow tube in the reactor. The pressure drop across the reactor can be as high as about 50 psi (3 atm) and thus the average pressure difference between the beds can be about 25 psi (1.5 atm).

The feed of synthesis gas 2 is introduced into the reactor via an inlet (not shown). The ratio of feed gas hydrogen to carbon monoxide is determined by the productivity and production of additional hydrogen using water gas conversion or the addition of hydrogen in addition to synthesis gas hydrogen into the reactor. It is generally high enough that carbon utilization is not negatively affected. The ratio of feed gas hydrogen to carbon monoxide is generally lower than the level at which excess methane will be produced. Advantageously, the ratio of hydrogen to carbon monoxide is between about 1.0 and about 2.2, and even between about 1.5 and about 2.2. If desired, pure synthesis gas can be used, and alternatively an inert diluent such as nitrogen, CO ₂ , methane, steam, etc. can be used. The term “inert diluent” indicates that the diluent is non-reactive under the reaction conditions or is a normal reaction product. It may be desirable to operate the synthesis gas conversion process in a partial conversion mode, for example 50-60% by weight, based on CO.

  According to one form of the method, the upstream bed 4 contains a Fischer-Tropsch synthesis gas conversion catalyst. The Fischer-Tropsch synthesis gas conversion catalyst can be cobalt, iron or ruthenium or a mixture comprising cobalt, iron or ruthenium. Preferred are catalysts having low water gas conversion activity suitable for reactions at lower temperatures, such as cobalt and ruthenium. The synthesis gas conversion catalyst can be supported on any suitable support such as, but not limited to, alumina, silica or titania or mixtures thereof. By way of non-limiting example, the synthesis gas conversion catalyst is supported in an amount between 5% and 50% by weight in the case of cobalt and between 0.01% and 1% by weight in the case of ruthenium. Can exist on top.

According to another form of the method, the upstream bed 4 contains a hybrid synthesis gas conversion catalyst. As used herein, the terms “hybrid Fischer-Tropsch catalyst” and “hybrid syngas conversion catalyst” convert the primary Fischer-Tropsch product to the desired product in a single stage (ie, heavier. Used interchangeably to denote a Fischer-Tropsch catalyst that includes a Fischer-Tropsch component as well as a component that contains the appropriate functionality to minimize the amount of unwanted product. The hybrid synthesis gas conversion catalyst contains a synthesis gas conversion catalyst in combination with an olefin isomerization catalyst, such as a relatively acidic zeolite, for isomerizing double bonds in C ₄₊ olefins as formed. To do. A method for preparing this type of hybrid catalyst is described in copending US patent application Ser. No. 12 / 343,534, which is hereby incorporated by reference in its entirety. Co-pending US patent application Ser. No. 12 / 343,534 uses a solution containing a cobalt salt to impregnate a zeolite extrudate to provide an impregnated zeolite extrudate, and by a reduction-oxidation-reduction cycle. Activating the impregnated zeolite extrudate is described. According to this method, the Fischer-Tropsch component (also referred to as “FT component” or “FT metal”) is not necessarily cobalt, but may be ruthenium, iron or cobalt, a mixture containing iron or ruthenium. Good. Impregnation of the zeolite using a substantially non-aqueous solution comprising a metal promoter salt and an FT metal salt, optionally activated later by a reduction-oxidation-reduction cycle, reduces ion exchange with the zeolitic acid sites. , Thereby increasing the overall activity of the zeolite component. The resulting catalyst comprises FT metal dispersed as small microcrystals on a zeolite support. The zeolite support, impregnation method and reduction-oxidation-reduction cycle used to activate the catalyst are described in detail below.

  The relatively larger zeolite extrudate particles result in lower pressure drops and experience less wear than zeolite powder or even granular zeolites (eg, having a particle size of about 300-1000 micrometers). The use of a zeolite extrudate as a support is beneficial. Methods for forming zeolite extrudates are readily known to those skilled in the art. A wide variety of macroporosities are possible with such extrudates. Without wishing to be bound by theory, as high a macroporosity as possible, consistent with a high enough crushing strength to allow operation in long reactor tubes, for hybrid synthesis gas conversion catalysts, selectivity and It is believed that it would be advantageous to minimize active diffusion constraints.

  The zeolite carrier is a molecular sieve containing silica at the tetrahedral framework position. Examples include silica only (silicate), silica-alumina (aluminosilicate), silica-boron (borosilicate), silica-germanium (germanosilicate), alumina-germanium, silica-gallium (gallosilicate) and silica-titania. (Titanosilicate), and mixtures thereof, including but not limited to.

  Molecular sieves are crystalline materials that in turn have regular passages (pores). If examined across several unit cells of the structure, the pores will form an axis based on the same unit in the repetitive crystalline structure. While the entire path of the pore is aligned with the pore axis in the unit cell, the pore may be separated from that axis, and it may be increased or decreased in size (forming the cage). May be. The axis of the pores is often parallel to one of the crystal axes. The narrowest position along the pore is the pore opening. The pore size indicates the size of the pore opening. The pore size is calculated by counting the number of tetrahedral positions that form the perimeter of the pore mouth. A pore having 10 tetrahedral positions at its pore mouth is commonly referred to as a 10-ring pore. The pores associated with catalysis in this application have a pore size of 8 or more rings. If the molecular sieve has the only type of associated pores with the same orientation axis as the crystal structure, it is called one-dimensional. The molecular sieve may have differently structured pores, or may have pores having the same structure but oriented in more than one axis associated with the crystal. In these cases, the dimensions of the molecular sieve are determined by summing the number of associated pores of the same structure but different axes and the number of associated pores of different shapes.

  Examples of zeolite carriers for hybrid synthesis gas conversion catalysts include SSZ-13, SSZ-33, SSZ-46, SSZ-53, SSZ-55, SSZ-57, SSZ-58, SSZ-59, SSZ-64, ZSM. Designed for -5, ZSM-11, ZSM-12, TS-1, MTT (eg SSZ-32, ZSM-23 etc.), HY, BEA (zeolite beta), SSZ-60 and SSZ-70 Including, but not limited to. These molecular sieves each contain silicon as the main tetrahedral element, have 8-12 ring pores, and are microporous molecular sieves, meaning that they have a pore opening of 20 rings or less. .

Zeolite support is between about 100 m ² / g to about 300 meters ² / g, it can have for example approximately 180 m ² / g, the external surface area of the. The micropore volume of 80% ZSM-5 is between about 90-112 μL / g with less volume suggesting loss of micropore structure or some blockage. The BET surface area is the sum of the micropore area and the outer area. The zeolite support further has a porosity between about 30-80%, a total penetration volume between about 0.25-0.60 cc / g, and a crush strength between about 1.25-5 lb / mm. be able to. The Si / Al ratio of the zeolite component can be between about 10-100.

  Initially, the zeolite support can be treated by oxidative calcination at a temperature in the range of about 450 ° C. to about 900 ° C., such as about 600 ° C. to about 750 ° C., to remove any organics and water from the zeolite support. .

  On the one hand, non-aqueous organic solvent solutions of FT component salts and optionally aqueous or non-aqueous organic solvent solutions of metal promoter salts are prepared, for example. Any suitable salt such as nitrate, chloride, acetate can be used. The aqueous solution for the promoter can be used in very small amounts. As used herein, the term “substantially non-aqueous” refers to a solution comprising at least 95% by volume non-aqueous solution. In general, any metal salt that is soluble in organic solvents and will not have a toxic effect on the catalyst may be utilized. The non-aqueous organic solvent is a non-acidic liquid formed from a site selected from the group consisting of carbon, oxygen, hydrogen, and nitrogen and has a relative volatility of at least 0.1. The term “relative volatility” refers to the ratio of the vapor pressure of acetone to the vapor pressure of acetone, measured at 25 ° C., as a reference. Suitable solvents include, for example, ketones such as acetone and butanone (methyl ethyl ketone); lower alcohols such as methanol, ethanol and propanol; amides such as dimethylformamide; amines such as butylamine; ethers such as diethyl ether and tetrahydrofuran; And a mixture of the aforementioned solvents. Suitable cobalt salts include, for example, cobalt nitrate, cobalt acetate, cobalt carbonyl, cobalt acetylacetonate, and the like. Similarly, any suitable ruthenium salt such as ruthenium nitrate, chloride, acetate can be used. In one form, ruthenium acetylacetonate is used. In general, any metal salt that is soluble in an organic solvent and would not have a toxic effect on the metal catalyst and on the acid sites of the zeolite can be utilized.

  The calcined zeolite support is then impregnated in a dehydrated state with a substantially non-aqueous organic solvent solution of the metal salt. Thus, the calcined zeolite support should not be excessively exposed to atmospheric humidity that causes rehydration. Any suitable impregnation technique can be used, including techniques well known to those skilled in the art, such as spreading the catalytic metal into a uniform thin layer on a catalytic zeolite support. For example, the FT component and the promoter can be deposited on the zeolite support material by the “initial wetness” technique. Such techniques are well known and are intended to predetermine a substantially non-aqueous solution volume to provide a minimum volume that will just wet the entire surface of the zeolite support without excess liquid. , Request. Alternatively, excess solution technology can be utilized if desired. If excess solution technology is utilized, then the excess solvent present, such as acetone, is simply removed by evaporation. Multiple impregnations are often required along with intermediate drying and calcination treatments to disperse and decompose the metal salts to achieve the desired metal formulation. The FT component content can vary from about 0.5% to about 25% by weight.

  A promoter metal may optionally be included in the hybrid synthesis gas conversion catalyst. For example, when the FT component is cobalt, suitable promoters include, for example, ruthenium, platinum, palladium, silver, gold, rhenium, manganese and copper. When the FT component is ruthenium, suitable promoters include, for example, rhenium, platinum, palladium, silver, gold, manganese and copper. By way of example, for a catalyst containing about 10 wt% cobalt, the amount of ruthenium promoter is about 0.01 to about 0.50 wt%, for example about 0.05 to about 0.25, based on the total catalyst weight. % By weight. The amount of ruthenium will be proportionally higher or lower, respectively, proportionally to higher or lower cobalt levels. A catalyst level of about 10% by weight has been found best for 80% by weight ZSM-5 and 20% by weight alumina. The amount of cobalt can be increased to about 20 wt% Co with increasing amount of alumina.

  The substantially non-aqueous solution and the zeolite support are then stirred while evaporating the solvent at a temperature of about 25 ° C. to about 50 ° C. until “dry”. The impregnated catalyst is gradually dried at a temperature of about 110 ° C. to about 120 ° C. for about 1 hour so that the metal spreads throughout the zeolite support. The drying process is performed in air at a very slow rate.

  The dried catalyst may be reduced directly in hydrogen or it may be calcined first. The dried catalyst is at a temperature in the range of about 200 ° C. to about 350 ° C., such as about 250 ° C. to about 300 ° C., sufficient to decompose the metal salt and immobilize the metal, for example 10 cc / gram / min. Baking is performed by gradually heating in an air stream. The drying and firing steps described above can be performed separately or together. However, the calcination should be carried out using a slow heating rate of, for example, 0.5 ° C. to about 3 ° C./min or about 0.5 ° C. to about 1 ° C./min, and the catalyst is about The maximum temperature should be maintained for 1 to about 20 hours, for example about 2 hours.

  The impregnation process described above is repeated with additional substantially non-aqueous solutions to obtain the desired metal formulation. Metal promoters can be added together with the FT component, but they may be added separately or in combination in other impregnation steps before, after, or during the impregnation of the FT component.

  After the last impregnation sequence, the blended catalyst zeolite support is then subjected to ROR activation treatment, which in turn is (A) reduction in hydrogen, (B) in oxygen-containing gas. Depending on the FT component used, the activation procedure is less than 500 ° C., even less than 450 ° C., even 400 ° C. Less than, or even less than 300 ° C. For the reduction step, temperatures between 100 ° C. and 450 ° C. and even between 250 ° C. and 400 ° C. are suitable. The oxidation step is between 200 ° C and 300 ° C. These activation steps are performed while heating at a rate of about 0.1 ° C. to about 5 ° C., such as about 0.1 ° C. to about 2 ° C. It has been found that when the catalyst is prepared by impregnating a zeolite support with an FT component such as cobalt or ruthenium, the activation procedure provides a catalyst with an improved reaction rate. Furthermore, when the promoter is pre-added, the activation procedure can significantly improve the activity of the catalyst.

  The ROR activation procedure of the present disclosure will now be described in more detail. The impregnated catalyst can be gradually reduced in the presence of hydrogen. If the catalyst is calcined after each impregnation to decompose nitrates or other salts, then the reduction is heated with a single temperature ramp (such as 1 ° C./min) to the maximum temperature. In an inert gas purge, performed in one step, at a temperature of about 250 ° C. or 300 ° C. to about 450 ° C., such as about 350 ° C. to about 400 ° C. It may be held with a holding time of 24 hours. Pure hydrogen is preferred in the first reduction step. If nitrate is still present, the reduction can be carried out in two steps, where the first reduction heating step is about 5 ° C./min or less, for example about 0.1 ° C. to about 1 ° C./min. Run at ambient pressure conditions at a slow heating rate up to a maximum holding temperature of 200 ° C. to about 300 ° C., for example 200 ° C. to about 250 ° C., for a holding time of about 6 to about 24 hours, eg about 16 to about 24 hours To do. In the second treatment step of the first reduction, the catalyst is about 0.5 ° C. to about 3 ° C./min, such as about 0.1 ° C. to about 1 ° C./min, about 250 ° C. or 300 ° C. to about 450 ° C. C., for example, to a maximum holding temperature of about 350.degree. C. to about 400.degree. C., for a holding time of 6 to about 65 hours, such as about 16 to about 24 hours. Pure hydrogen is preferred for these reduction steps, but a mixture of hydrogen and nitrogen can be utilized.

  The reduction uses a mixture of hydrogen and nitrogen at 100 ° C. for about 1 hour; increasing the temperature to 0.5 ° C./min to a temperature of 200 ° C .; holding at that temperature for about 30 minutes; and then 350 ° C. The reduction may be continued for about 16 hours at a temperature of 1 ° C./min until a temperature of The reduction should be carried out gradually and sufficiently, maintaining a flow of reducing gas high enough to maintain a partial pressure of water in the off-gas of less than 1%. Before and after all reductions, the catalyst is purged in an inert gas such as nitrogen, argon or helium.

  The reduced catalyst is passivated at ambient temperature (25 ° C. to 35 ° C.) by flowing dilute air over the catalyst sufficiently slowly so that a controlled exotherm of + 50 ° C. or less passes through the catalyst bed. Is done. After passivation, the catalyst is gradually heated in diluted air to a temperature of about 300 ° C. to about 350 ° C. (preferably 300 ° C.) in the same manner as described above in connection with the calcination of the catalyst. .

  The reoxidized catalyst is then gradually reduced again in the presence of hydrogen in the same manner as described above in connection with the initial reduction of the impregnated catalyst. This reduction may be achieved and maintained with a single temperature ramp as described above for the reduction of the calcined catalyst.

  While the ROR activation procedure of the present disclosure may be used to improve the activity of the hybrid syngas conversion catalyst of the present disclosure, those skilled in the art to spread the catalyst metal in a uniform manner on the catalyst zeolite support. Any known technique is suitable, although they do not promote ion exchange with the zeolitic acid sites.

  The hybrid syngas conversion catalyst has an average particle size of about 0.01 to about 6 millimeters, such as about 1 to about 6 millimeters.

  According to yet another aspect, the upstream bed 4 contains a mixture of a conventional Fischer-Tropsch catalyst and a hybrid synthesis gas conversion catalyst, and the bed is about 1 to about 99 wt% based on the total catalyst weight. Conventional Fischer-Tropsch catalyst and between about 1 and about 99% by weight of hybrid syngas conversion catalyst.

  The intermediate catalyst bed 6 contains a hydroisomerization catalyst for hydroisomerizing linear hydrocarbons. The hydroisomerization catalyst is a bifunctional catalyst containing a hydrogenation component including an acidic component and a metal promoter. The hydroisomerization catalyst can be selected from 10 or more zeolites. Suitable materials for use as hydroisomerization catalysts include, but are not limited to, SSZ-32, ZSM-57, ZSM-48, ZSM-22, ZSM-23, SAPO-11 and theta-1 examples. Not. The hydroisomerization catalyst can also be a non-zeolitic material.

  According to one form, the intermediate catalyst bed 6 also contains a hydrocracking catalyst for cracking linear hydrocarbons. The hydrocracking catalyst is an acid catalyst material and can be a material such as amorphous silica-alumina or tungstated zirconia or a zeolitic or non-zeolitic crystalline molecular sieve. The crystallographic free diameter of the molecular sieve channel is described in “Atlas of Zeolite Framework Types”, Fifth Revised Edition, 2001, by CH. Baerlocher, W.M. M.M. Meier, and D.M. H. Olson. Published in Elsevier, pp 10-15, incorporated herein by reference. Examples of suitable hydrocracking intermediate pore molecular sieves include zeolite Y, zeolite X and so-called ultrastable zeolite Y and high structure silica: alumina ratio zeolite Y, for example, US Pat. 401, 556, 4, 820, 402 and 5,059, 567. The term “intermediate pore” as used herein means having a crystallographic free diameter in the range of about 3.9 to about 7.1 angstroms when the molecular sieve is in the fired shape. Small crystal size zeolite Y can also be used, such as those described in US Pat. No. 5,073,530, incorporated herein by reference. Other zeolites showing utility as cracking catalysts are SSZ-13, SSZ-33, SSZ-46, SSZ-53, SSZ-55, SSZ-57, SSZ-58, SSZ-59, SSZ-64, ZSM. -5, ZSM-11, ZSM-12, ZSM-23, HY, beta, mordenite, SSZ-74, ZSM-48, TON type zeolite, ferrierite, SSZ-60 and SSZ-70 including. Non-zeolitic molecular sieves that can be used include, for example, silicoaluminophosphates (SAPO) such as SAPO-11, SAPO-31 and SAPO-41, ferroaluminophosphates, titanium aluminophosphates and US Pat. 913,799 and the various ELAPO molecular sieves described in the references cited therein. Details regarding the preparation of various non-zeolitic molecular sieves can be found in various references cited in US Pat. No. 5,114,563 (SAPO), US Pat. No. 4,913,799, and US Pat. No. 4,913,799. And can be found in their entirety here by reference. The molecular sieves of mesopores are, for example, M41S-based materials (J. Am. Chem. Soc. 1992, 114, 10834-10843), MCM-41 (US Pat. Nos. 5,246,689, 5,198,203, 5,334,368), and MCM48 (Kresge et al., Nature 359 (1992) 710).

  The hydrocracking and hydroisomerization catalyst can optionally contain a metal promoter and cracking components. The metal promoter is typically a metal or combination of metals selected from Group VIII noble and non-noble metals, Group 1B cast coin metals, and Group VIB metals. Noble metals and cast coin metals that can be used include platinum, palladium, rhodium, ruthenium, osmium, silver, gold and iridium, or any combination thereof. Non-noble metals that may be used include molybdenum, tungsten, nickel, cobalt, copper, rhenium, or any combination thereof.

  The metal promoter can be incorporated into the catalyst mixture by any one of numerous procedures. It can be added to either the degradation component, the carrier, or a combination of both. Alternatively, the Group VIII component can be added to the decomposition or matrix component by co-polishing, impregnation, or ion exchange, and the Group VI components, ie, molybdenum and tungsten, can be fire-oxidized by impregnation, co-polishing or co-precipitation. Can be combined with things. These components are usually added as metal salts that can be reduced to metals with hydrogen or other reducing agents or thermally converted to the corresponding oxides in an oxidizing atmosphere.

According to one form, the intermediate catalyst bed 6 contains a combination of a solid acid hydrocracking component such as Pd / ZSM-5 and a hydroisomerization component such as SSZ-32 based noble metal promoted zeolite. The ratio of cracking catalyst and hydroisomerization catalyst in the intermediate bed is advantageously optimized to balance the isomerization activity with the cracking activity. If excessive cracking catalyst is present, the resulting product may be lighter than desired. While the cracking catalyst converts the n-paraffin wax product to the appropriate chain length, the hydroisomerization component isomerizes the n-paraffin product into a completely liquid isomerization product. If it is desired to produce heavier diesel range products, then the catalyst combination should exhibit less cracking and more isomerization. For example, it has been found that more isomerization can be achieved by including Pd / SSZ-32. If insufficient cracking catalyst is present, the hydroisomerization catalyst may not be able to convert the wax to a liquid product under the mild process conditions of the process. Thus, for example, to obtain the desired product having an average molecular weight in the diesel range, ie C ₁₁ -C ₂₀ , and not containing a solid wax phase at ambient conditions, the correct proportions of cracking catalyst components and hydroisomerization It may be advantageous to include a combination of both catalysts in the intermediate bed.

  The amount of catalyst mixture required in the intermediate bed will depend in part on the tendency of the syngas conversion catalyst to produce wax in the upstream bed and in part on the process conditions. Generally, the weight of the catalyst mixture in the intermediate bed is between about 0.5 to about 2.5 times the weight of the catalyst in the upstream bed.

  The reaction temperature is suitably about 160 ° C to about 260 ° C, such as about 175 ° C to about 250 ° C or about 185 ° C to about 235 ° C. Higher reaction temperatures favor lighter products. The total pressure is, for example, from about 1 to about 100 atmospheres, such as from about 3 to about 35 atmospheres, or from about 5 to about 20 atmospheres. Higher reaction pressure favors heavier products. The gas hourly space velocity based on the total amount of feed is less than 20,000 gas volume per hour of catalyst volume, for example about 100 to about 5000 v / v / hour or about 1000 to about 2500 v / v / hour.

  Fixed bed reactor systems have been developed to perform Fischer-Tropsch reactions. Such a reactor is suitable for use in the present method. For example, a suitable Fischer-Tropsch reactor system includes a multi-tube fixed bed reactor in which the catalyst is compounded into the tubes.

The process in the upstream and intermediate beds provides high yields of middle distillate and / or light base oil range paraffinic hydrocarbons under essentially common reaction conditions. The produced intermediate mixture 3 is liquid at about 0 ° C. Mixture 3 is substantially free of solid wax, which means that the product is a single liquid phase at ambient conditions without the visible turbid presence of an insoluble solid wax phase. “Ambient conditions” means a temperature of 15 ° C. and a pressure of 1 atmosphere. The intermediate mixture 3 has the following composition:
0-20, for example, 5-15 or 8-12 wt% of _CH 4;
0-20, for example, of 5-15 or 8-12 wt% _C 2 _-C 4;
60-95, for example, 70-90 or 76-84% by weight of _{C 5+;} and 0-5 wt% of _{C 21+} normal paraffins.

In a typical Fischer-Tropsch process, the resulting product is primarily a normal or linear paraffin product, which means no branching. If more than 5% by weight of C ₂₁₊ fraction present mainly inside the linear product, the product was found to contain a separate visible solid wax phase. The product of the process may actually contain more than 5% by weight of C ₂₁₊ without a visible solid wax phase. This appears to be due to the hydroisomerization capability of the hydroisomerization catalyst. Branched paraffins have a lower melting point compared to normal or linear paraffins, so that the product of the process can contain a higher percentage of C ₂₁₊ fraction and still be separated at ambient conditions. It can remain liquid without any visible solid wax phase. The method provides a product having a C ₂₁₊ paraffin concentration (determined by gas chromatography) of at least 30% by weight isomerization (ie containing at least a single branch) based on the weight of the C ₂₁₊ fraction. . The result is a liquid product that can be poured at ambient conditions. Liquid hydrocarbons produced by this process advantageously have a low cloud point of 15 ° C. or lower, further 10 ° C. or lower, further 5 ° C. or lower, and even 2 ° C., as determined by ASTM D 2500-09. .

In addition, the process eliminates the need for separation of the product resulting from the first catalyst bed and intermediate distillation without the need for a second reactor containing a hydrocracking and / or hydroisomerization catalyst. Products and / or paraffinic hydrocarbons in the light base oil range in high yield. The process water originating from the first catalyst bed is not required to be separated from the reactor during the hydroisomerization of the C _{21 +} normal paraffin. With a suitable combination of catalyst composition, catalyst bed arrangement and reaction conditions, both the syngas conversion reaction and the subsequent hydrocracking and / or hydroisomerization reaction can be performed in a single reactor under essentially common process conditions. It was found that it can be implemented within.

  The intermediate mixture 3 is optionally sent to a separator 15 that utilizes a drop in temperature to condense the water 17 and separate the intermediate mixture 21 and the gas stream 19. The gas stream 19 is recycled to the upstream bed via a compressor (not shown). In order to reduce the content of unconverted gas recycled to the upstream bed, a part of the gas stream 19 is fed into a synthesis gas generation unit such as a flare (not shown) or an autothermal reformer (not shown). Is sent in some cases.

  The intermediate mixture 3 (or optionally the intermediate mixture 21 when using the separator 15) is passed through a downstream catalyst bed 16 containing an olefin saturated catalyst capable of saturating the olefins contained in the intermediate mixture. A hydrogen stream 28 is fed to the catalyst bed 16. The resulting product stream 30 contains no more than about 25 weight percent olefins, preferably no more than about 5 weight percent olefins, and is essentially free of olefins. “Essentially free of olefins” means that the product stream has a bromine number of less than about 1.

  According to this form, the downstream catalyst bed 16 can contain a catalyst containing a hydrogenation component useful as an olefin saturation catalyst deposited on a support. The hydrogenation component can be a Group IB noble metal, a Group VIII noble metal, or a combination thereof. Preferred noble metals include platinum, palladium, rhodium, iridium, silver, osmium and gold, and combinations thereof, and combinations of metals including ruthenium. According to this configuration, the downstream catalyst bed 16 can also include an olefin saturation catalyst having a hydrogenation component selected from a Group VIII non-noble metal and a Group VIB metal or combination. Non-noble metals that can be used include molybdenum, tungsten, nickel, iron, zinc, copper, lead and cobalt. Suitable combinations of metals include at least one Group VIII metal and one Group VIB metal, such as nickel-molybdenum, cobalt-molybdenum, nickel-tungsten, and cobalt-tungsten. The overall catalyst composition of preferred non-noble metals is at least about 0.5, and generally from about 1 to about 15 weight percent nickel and / or cobalt, and greater than about 5 weight percent, as determined as the corresponding oxide, Preferably from about 5 to about 40 weight percent molybdenum and / or tungsten. The non-noble metal hydride metal can be present in the olefin saturated catalyst composition as a metal sulfide when the compound is readily formed from the particular metal involved. These metal sulfide forms may have desirable activity, selectivity and activity retention.

  The olefin saturation catalyst can be supported on any suitable support such as, but not limited to, alumina, silica or titania, or mixtures thereof, such as solid oxides. This support may be a zeolite support containing silica in the tetrahedral framework position. Examples include silica only such as silicalite (silicate), silica-alumina (aluminosilicate), silica-boron (borosilicate), silica-germanium (germanosilicate), aluminum-germanium, silica-gallium (gallosilicate). And silica-titania (titanosilicate), and mixtures thereof.

According to this configuration, the pressure in the downstream catalyst bed 16 is between about 200 psig (1.4 MPa) and about 3000 psig (21 MPa), preferably between about 500 psig (3.4 MPa) and about 2000 psig (13 MPa). . The temperature range in the downstream catalyst bed 16 is typically between about 300 ° F. (150 ° C.) and about 700 ° F. (370 ° C.), preferably about 400 ° F. (205 ° C.) to about 500 ° F. (260 ° C.). ). LHSV is usually in the range of about 0.2 to about 2.0 h ⁻¹ , preferably 0.2 to 1.5 h ⁻¹ and most preferably about 0.7 to 1.0 h ⁻¹ . Hydrogen at a rate of feed per barrel to about 1000SCF ^(28m 3) ~ about 10,000SCF (280m ^3), usually in the downstream catalyst bed 16 is supplied. Typically, hydrogen is supplied at a rate of about 3000 SCF (85 m ³ ) per feed barrel.

  According to another form, the upstream, intermediate and downstream beds are physically arranged in series in the same reactor. FIG. 2 illustrates this configuration, where the syngas feed 2 enters the reactor and the upstream catalyst bed 4 of the syngas conversion catalyst, the intermediate catalyst bed 6 of the hydroisomerization catalyst, and the olefin saturation catalyst. The synthesis gas conversion catalyst, hydroisomerization catalyst, and olefin saturation catalyst are each as described above. Stream 18 optionally passes through separator 14 and stream 12 is optionally recycled.

  According to yet another aspect, the hydroisomerization catalyst and olefin saturation catalyst are physically mixed in the same bed 10 in the same reactor downstream of the catalyst bed 4. FIG. 3 illustrates this configuration, where the syngas feed 2 enters the reactor and of the upstream catalyst bed 4 of the syngas conversion catalyst, the hydroisomerization catalyst and the downstream catalyst bed 10 of the olefin saturation catalyst. They are as described above. The downstream bed 10 can contain between about 10 to about 90 wt% hydroisomerization catalyst and about 90 to about 10 wt% olefin saturation catalyst, respectively, based on the total catalyst weight in the bed 10. Stream 18 optionally passes through separator 14 and stream 12 is optionally recycled.

Examples In the following examples and comparative examples, a 5 mm ID, 300 mm long stainless steel tube reactor was used. The reactor included a 140 mm long heating zone. The catalyst was held at 80-100 mm in the center of the heating zone.

  The catalyst was subjected to the following ROR activation procedure. Initial reduction in 96% H2-4% Ar at 10 atm by ramping from ambient temperature to 350 ° C. at 0.5 ° C./min and holding at intermediate temperatures of 150 ° C. and 250 ° C. for 3 hours respectively. And the reduction was held at 350 ° C. for 12 hours. The flow was then switched to 96% N2-4% Ar for purging and the temperature was reduced to 100 ° C. over 12-15 hours and held at 100 ° C. for an additional 5 hours. The oxidation was then performed by switching to a flow of 5% O2-91% N2-4% Ar, heating to 300 ° C. at 1 ° C./min, and then holding at that temperature for an additional 6 hours. The stream was returned to 96% N2-4% Ar and the temperature was reduced to 100 ° C. over 10-12 hours. The catalyst was held at that temperature for an additional 16 hours. Final reduction was performed in 96% H2-4% Ar with 10 atm and 62.5 sccm flow. The heating lamp was 0.25 ° C./min in this process. The final holding temperature was 250 ° C. After holding at 250 ° C. for 12 hours, the temperature was reduced to about 150 ° C. for the start of the test run.

  The test was started at 150 ° C. in a flow of 64% H2-32% CO-4% Ar (H2 / CO = 2). The pressure was 10 atm and the flow was 2.4 NL / h per position. The temperature was raised to 180 ° C. over the first hours of operation. Thereafter, when the temperature was rising, the temperature was never increased faster than 0.01 ° C./min. Temperature changed from 180 ° C to 230 ° C, H2 / CO ratio from 1.5 to 2.0, pressure from 10 to 20 atm, and flow rate from 1.2 to 2.4 NL / reactor / h. did. H2, CO, CO2, and total hydrocarbons were analyzed online by an ABB analyzer. C1-C12 hydrocarbons were analyzed online by gas chromatography (FID and TCD based); sufficient isomer distribution was possible through C5. C13 + hydrocarbons were collected in a hot trap and then analyzed offline on a high temperature GC instrument designed for wax analysis. If the product exiting C65 was present, it was analyzed. The bromine number was determined by the ASTM D1159 test.

Example 1
150 mg synthesis gas conversion (FT) catalyst comprising 0.19 wt% Ru and 7.5 wt% Co supported on 20 wt% alumina and 80 wt% ZSM-12 was prepared as follows: ˜160 μm in size and diluted to 4 volumes with carborundum. The FT catalyst was placed in a reactor upstream of a mixture of 253 mg Pd / SSZ-32 and 237 mg Pt / ZSM-5 prepared as follows. The total length of both floors was 75 mm. The corresponding volume was 1.47 mL.

Preparation of FT catalyst The catalyst was prepared using a three-step initial wetness impregnation method. 125.824 g of cobalt (II) nitrate hexahydrate (obtained from Sigma-Aldrich), 2.041 g of ruthenium (III) nitrosyl nitrate (obtained from Alfa Aesar) and (obtained from Sigma-Aldrich) in water ) A solution was prepared by dissolving 3.381 g of lanthanum (III) nitrate hexahydrate. After calcining at 750 ° C. for 2 hours in air, 100 g of Puralox® alumina SBA200 support (obtained from Sasol) was impregnated by using one third of this solution to achieve initial wetness did. The prepared catalyst was then dried in air in a box furnace at 120 ° C. for 16 hours, and then the temperature was increased to 300 ° C. at a heating rate of 1 ° C./min and held at that temperature for 2 hours. It was then fired in air by cooling back to ambient temperature. The above procedure was repeated to obtain the following formulation of Co, Ru and La ₂ O ₃ on a carrier with a formulation of 20 wt% Co, 0.5% Ru and 1 wt% La ₂ O ₃ and 78.5% alumina. It was.

Preparation of Pt / ZSM-5 catalyst (obtained from Aldrich) 1.2466 g of tetraammineplatinum (II) nitrate was dissolved in 60 cc of water. The resulting solution was added to 125.1 g of CBV8014 zeolite extrudate (obtained from Zeolyst International). Most of the water was removed in a rotary evaporator under vacuum by heating slowly to 65 ° C. The vacuum dried material was then further dried in an oven at 120 ° C. overnight. The dried catalyst was calcined at 300 ° C. for 2 hours in a muffle furnace.

The Pd / SSZ-32 Preparation of catalyst (obtained from Aldrich) _{H 2} O in 10 wt% of 28.34g tetraamine palladium (II) nitrate solution were dissolved in water of 100 cc. The resulting solution was added to Chevron USA Inc. To 100 g of commercially available SSZ-32 based zeolite. Most of the water was removed in a rotary evaporator under vacuum by gradually heating to 65 ° C. The vacuum dried material was then further dried in an oven at 120 ° C. overnight. The dried catalyst was calcined at 300 ° C. for 2 hours in a muffle furnace.

Comparative Example 1
150 mg of CoRu / alumina catalyst was diluted to 4 volumes as described in Example 1 and placed in the upstream bed. The total length of the catalyst bed was 70 mm. The corresponding volume was 1.37 mL.

Comparative Example 2
150 mg of CoRu / alumina catalyst was diluted to 4 volumes as described in Example 1 and placed in the upstream bed. The Pt / ZSM-5 catalyst described in Example 1 was placed in the 473 mg downstream bed. The total length of the upstream catalyst bed and the downstream catalyst bed was 70 mm. The corresponding volume was 1.37 mL.

  Table 1 shows the results including three runs for each example and comparative example. The results illustrate the effectiveness of the present method to significantly reduce the olefin content of the intermediate hydrocarbon mixture.

Claims (23)

A feed comprising a mixture of hydrogen and carbon monoxide is fed into the following a. Synthesis gas conversion catalyst, b. A hydroisomerization catalyst, and c. A process for converting synthesis gas to a hydrocarbon mixture comprising contacting an olefin saturated catalyst in sequence with:
a. A synthesis gas conversion catalyst in the upstream bed, wherein a first intermediate hydrocarbon mixture containing C _{21 +} normal paraffin and olefin is formed on the synthesis gas conversion catalyst;
b. A hydroisomerization catalyst containing an acidic component and a metal promoter downstream of the intermediate catalyst bed of the upstream catalyst bed, wherein the C _{21 +} normal paraffin of the first intermediate hydrocarbon mixture is a hydroisomerization catalyst. _{Hydroisomerized above} , thus forming a second intermediate hydrocarbon mixture containing up to about 5% by weight of _{C21 +} normal paraffins and olefins, and c. An olefin saturation catalyst downstream of the intermediate catalyst bed downstream of the intermediate catalyst bed, wherein the olefin is saturated on the olefin saturation catalyst;
Thereby resulting in a final hydrocarbon mixture containing no more than about 5% by weight C 21 ₊ normal paraffins and no more than about 25% by weight olefins.
Said method.   The process of claim 1, wherein the final hydrocarbon mixture contains up to about 5 wt% olefins.   The process of claim 1 wherein the final hydrocarbon mixture is essentially free of olefins.   2. The upstream bed and the intermediate bed are in a single multitubular fixed bed reactor and have an essentially common reactor temperature and an essentially common reactor pressure. the method of.   The process according to claim 1, wherein the hydroisomerization catalyst comprises SSZ-32 based zeolite.   The method of claim 1, wherein the intermediate bed further comprises a hydrocracking catalyst selected from the group consisting of amorphous silica-alumina, tungstated zirconia, zeolitic molecular sieves and non-zeolitic crystalline molecular sieves. The process of claim 4, wherein process water is not separated from the reactor during the hydroisomerization of the C _{21 +} normal paraffin.   5. The method of claim 4, wherein no hydrogen is added to the reactor in addition to the hydrogen and carbon monoxide mixture.   The method of claim 1, wherein the hydrocarbon mixture is substantially free of solid wax at ambient conditions. The process according to claim 1, wherein the hydrocarbon mixture has an isomerized _{C21 +} paraffin concentration of at least 30% by weight, based on the weight of the _{C21 +} fraction.   The method of claim 1, wherein the hydrocarbon mixture has a cloud point of 15 ° C. or less.   Between the intermediate catalyst bed and the downstream catalyst bed, a second intermediate hydrocarbon mixture passes through a separator and is recycled to the upstream catalyst bed, water removed and the downstream catalyst. The process of claim 1, wherein the process is separated into liquid hydrocarbons that pass through the bed.   The process of claim 1, wherein the olefin saturation catalyst is selected from the group consisting of metals selected from Group IB noble metals and Group VIII noble metals and combinations thereof.   The method of claim 1, wherein the olefin saturation catalyst is selected from the group consisting of platinum, palladium, rhodium, iridium, silver, osmium and gold, and combinations thereof.   13. The method of claim 12, wherein the olefin saturation catalyst is selected from the group consisting of metals selected from Group VIII noble and non-noble metals and Group VIB metals and combinations thereof.   13. The olefin saturation catalyst is selected from the group consisting of molybdenum, tungsten, nickel, iron, zinc, copper, lead, cobalt, nickel-molybdenum, cobalt-molybdenum, nickel-tungsten, and cobalt-tungsten. the method of.   The method of claim 12, wherein the olefin saturation catalyst comprises a metal sulfide.   The process of claim 1, wherein the olefin saturation catalyst comprises a zeolite support.   The process of claim 1, wherein the olefin saturation catalyst comprises a zeolite comprising SSZ-32.   The process of claim 1, wherein the upstream catalyst bed temperature is between about 160 ° C. and about 260 ° C.   The method of claim 4, wherein the temperature of the upstream catalyst bed and the temperature of the intermediate catalyst bed differ by about 20 ° C. or less. The final hydrocarbon mixture produced is:
0-20 wt% of _CH 4;
0-20 wt% _C 2 -C _4; and 60 to 95% by weight of _{C 5+}
The method of claim 1 comprising: A feed comprising a mixture of hydrogen and carbon monoxide is fed into a. A synthesis gas conversion catalyst, and b. A process for converting synthesis gas to a hydrocarbon mixture comprising sequentially contacting a mixture of a hydroisomerization catalyst and an olefin saturation catalyst, comprising:
a. A synthesis gas conversion catalyst in the upstream bed, wherein a first intermediate hydrocarbon mixture containing C _{21 +} normal paraffin and olefin is formed on the synthesis gas conversion catalyst;
b. A mixture of a hydroisomerization catalyst and an olefin saturation catalyst containing an acidic component and a metal promoter in the downstream catalyst bed, whereby the C _{21 +} normal paraffin of the intermediate hydrocarbon mixture is hydroisomerized; And the olefin is saturated, thus forming a hydrocarbon mixture containing no more than about 5% by weight C 21 ₊ normal paraffins and no more than about 25% by weight olefins.
Said method.