MXPA00002899A - A process for optimizing hydrocarbon synthesis - Google Patents

Abstract

The invention is a process for producing hydrocarbons from hydrogen and carbon monoxide by reacting hydrogen and carbon monoxide in the presence of a particulate solid catalyst and a substantially inert liquid medium. This reaction takes place in a reactor vessel adapted for the reaction of gases in the presence of a substantially inert liquid medium and a bed of solid particulate catalyst. The hydrogen gas and carbon monoxide gas are introduced at a plurality of locations within the reactor vessel. Bubbles of gas flow upward through the bed of solid catalyst particles and substantially inert liquid medium at sufficient velocity toexpand the bed to a volume greater than its static volume. This velocity creates a turbulent reaction zone wherein liquid, gas, and solid catalyst are present and are in a state of motion.

Description

A PROCESS TO OPTIMIZE IA SYNTHESIS OF HYDROCARBON FIELD OF THE INVENTION This invention relates to the production of hydrocarbons from hydrogen gas and carbon monoxide gas. Specifically, the invention is. relates to the production of hydrocarbon in a boiling bed.

BACKGROUND OF THE INVENTION The reaction to convert mixtures of carbon monoxide and hydrogen (defined here as synthesis gas) to hydrocarbons on metal catalysts has been known since the beginning of the century. This reaction is commonly known as the Fischer-Tropsch or F-T synthesis. The synthesis gas used as food for the process can be obtained from any source known to those skilled in the art, such as, for example, steam reforming of natural gas or partial oxidation of coal. An important criterion for commercial F-T synthesis is having the ability to control the temperature of the reagents. The F-T reaction is highly exothermic. The efficient and rapid removal of heat is the main consideration in the generation of high molecular weight hydrocarbons. Unfortunately, high temperatures, ie, above 325 ° C, often lead to methane generation, carbon deposition on the catalyst, and fragmentation of catalyst particles. Methane generation is usually not desirable because the yield of the higher hydrocarbons is reduced. The deposition of carbon and fragmentation of catalyst particles is undesirable because the life of the catalyst is shortened. The prior art solves the problem of heat generation of the highly exothermic F-T reaction using long tubular reactors that have a greater surface area to volume ratio than more conventional cylindrical reactors, thus using the additional surface area for cooling. Another method used by the prior art is to effect the reaction at low conversion rates by passing through the reactor, thereby using the unreacted gas to remove heat. The temperature gains can also be controlled using a suspension bed reactor. The main disadvantages of the commercialization of the suspension bed reactor process in the prior art are the separation of serous products and fine particles of catalyst, and the mechanical failure due to the high erosion of the pumping equipment used to recycle the suspension to the reactor area.

Another problem with the commercialization of F-T is the efficient conversion of the reagents. F-T synthesis generally uses hydrogen and carbon monoxide at a molar ratio just above 2.0: 1. stechymetrically, a hydrogen molecule combines with carbon to form hydrocarbons and a second hydrogen molecule combines with oxygen to form water vapor. The gas in the reactor can be depleted in a reagent, which will slow the reaction rate to levels below commercial viability. The reduction in the reaction rate is exacerbated when the reagent is deficient in hydrogen. There are many different catalysts available for an F-T synthesis, and the effect of a deficiency on a reactant on the reaction rate varies between those catalysts. However, the effect of the 3s hydrogen gas deficiency is always more pronounced than the effect of a carbon monoxide deficiency. For example, the reaction rate F-T with a cobalt-based catalyst is increased with the partial pressure of hydrogen and carbon monoxide; However, changes in the partial pressure of hydrogen have almost twice the effect that changes in the partial pressure of carbon monoxide. The reaction rate therefore falls twice as fast when the synthesis gas is deficient in hydrogen compared to the decline in the reaction rate caused by a deficiency in carbon monoxide. This is usually not problematic if the ratio of hydrogen to carbon monoxide in the synthesis gas is about 2.1: 1 and there is little methane synthesis. Nevertheless, to obtain a greater fraction of waxy hydrocarbon product, a ratio of hydrogen to carbon monoxide less than 2.1: 2 may be required. therefore, as an F-T synthesis of the reaction of the waxy hydrocarbons proceeds and the synthesis gas is converted to hydrocarbons, the synthesis gas is progressively depleted in hydrogen. This results in a substantial portion of the hydrogen gas, carbon monoxide gas or both, leaving the reactor without being converted. Another related problem is that the reactive gases are converted into hydrocarbons and water, gases diluted in the feed gas stream, for example water vapor, light hydrocarbons and contaminants, which can dilute the hydrogen gas and the carbon monoxide gas to the point at which the reaction rate is significantly reduced. This further exacerbates the reduction in speed that is experienced when the reagent becomes deficient.

Finally, in typical reactors where a reactant is gas, the distribution of gas in the reaction zone and the back-mixing are main factors that determine the operation of the reactor. Poor gas distribution will result in a stagnant flow in suspension or channeling reactors in tubular reactors which results in the reactive gases not being uniformly exposed to the catalyst. Poor distribution of the gas results in a hot spot in the reactor which favors the undesirable methane formation reaction and can damage the catalyst. Backmixing will reduce kinetic performance. Such maldistribution and back-mixing often occur in conventional F-T processes. It would be desirable if a process could be developed F-T that will provide sufficient temperature control to prevent methane formation, deactivation of the catalyst through the deposition of carbon, and fragmentation of the catalyst. It would also be desirable if such a process offered a way to maintain a hydrogen to carbon monoxide ratio different from the stoichiometric ratio of 2.1: 1. It would be further desirable that such a process provide means for distributing the reactive gases in the catalyst bed in such a way that the channeling or backmixing would be reduced.

BRIEF DESCRIPTION OF THE INVENTION The invention is a process for producing hydrocarbons from hydrogen without carbon monoxide by reacting hydrogen and carbon monoxide in the presence of a solid particulate catalyst. This reaction takes place in a reactor vessel adapted for the gas reaction in the presence of a substantially inert liquid medium and a bed of solid particulate catalyst, ie a boiling bed reactor. The hydrogen gas, the carbon monoxide gas, or both are introduced into a plurality of places within the reactor vessel. The liquid medium is introduced at least in one place in the reactor vessel and is generally well mixed with the gas before entering the reaction zone. Bubbles of gas and substantially inert liquid medium flow through the bed of solid catalyst particles at a rate sufficient to expand the reaction zone to a volume greater than its static volume, this velocity creates a turbulent reaction zone where the Liquid, gas and solid catalyst are present in a constant state of motion. The process employed in the present invention is often described as a boiling bed process. The boiling bed process comprises the passage of liquids and gas flows flowing substantially concurrently through a vertical cylindrical vessel containing catalyst particles. The catalyst is moving in the fluid comprised of liquid and gas bubbles. The movement of the catalyst bed is controlled by the gas flow and by the liquid flow, so that in a steady state the catalyst volume is controlled to a definable level in the reactor. The boiling bed process of the present invention provides multiple options for temperature control and catalyst level control, while maintaining very high throughput and conversion, heat is removed in a variety of ways, including the control of the gas fed preheated to a desired temperature, internal heat exchangers, removal of a flow of liquid medium and product for cooling and return of cooled liquid medium or addition of liquid medium from another source to the. reactor. In addition, the sensible and latent heat capacity of the liquid medium moderates the temperature changes. The reactor fluid provides convenient means for removing product from the reactor. In addition, the movement of the catalyst in the reactor provides convenient means to remove the catalyst while the reactor is in flow. The rate of catalyst addition and extraction is used to control the activity of the catalyst. The introduction of gas to a plurality of places within the reactor provides means to supplement the gas within the reactor with hydrogen gas, carbon monoxide gas or other gases, or mixtures thereof, at predetermined points in the reactor. This supplemental gas is used to maintain the optimum ratio of hydrogen to carbon monoxide, so that the selectivity of the catalyst is toward the production of the desired product, which in a preferred embodiment of this invention is a serous paraffin. This supplemental gas is used to maintain the partial pressure of both carbon monoxide and hydrogen to facilitate a high but controlled reaction rate. In addition, this injection of supplemental gas will provide the gas distribution through the event and will help prevent undesirable backmixing.

BRIEF DESCRIPTION OF THE DRAWINGS Figure 1 is a schematic diagram of a reactor in the boiling bed for the conversion of synthesis gas to liquid hydrocarbons. A commercial modality of the F-T synthesis process can be optimized to produce mainly fuels and lubricants as products. Figure 2 is a schematic of such a process where the boiling bed reactor 1 is designed for a very high conversion to a single performance. The internal parts of the reactor are similar to those shown in Figure 1 and are not shown here. Another commercial application of the F-T synthesis process produces mainly fuels as products. Figure 3 is a schematic of such a process where three reactors are used in series to process the fed synthesis gas. In this scheme, the F-T synthesis occurs in the first two reactors. The function of the third reactor is not that of an F-T synthesis but that of decomposing hydrocarbons with high molecular weight in low molecular weight hydrocarbons.

DETAILED DESCRIPTION OF THE INVENTION As used herein, the term "reactive gas" refers to any carbon monoxide gas or hydrogen gas, or a mixture of both. The mixture comprising these gases is often referred to as "synthesis gas" or "gas without". A synthesis gas can also contain inert and polluting compounds. Examples of inert compounds include carbon dioxide, nitrogen gas and methane. Carbon dioxide and methane are usually not desired feed components for use with the process of this invention, but in smaller amounts usually do not adversely affect the process, unless they act as diluents. Examples of contaminants include molecular oxygen, metallic carbonyl compounds, water vapor and sulfur-containing compounds. The presence of water vapor, a product of the F-T reactions, in the feed gas will slightly reduce the equilibrium conversion of reactive gases, but this smaller amount has little practical effect in a commercial reactor where large amounts of water are generated. The compounds they contain and the metal carbonyl compounds, on the other hand, can contaminate the catalyst and should be kept to a minimum. As used herein the term "hydrocarbon" is used generically and refers to compounds containing hydrogen and carbon. These include straight chain hydrocarbons, cyclic hydrocarbons, branched hydrocarbons, saturated or paraffinic hydrocarbons, unsaturated hydrocarbons, olefins, oxygenates and other compounds containing predominantly hydrogen and carbon. Paraffinic hydrocarbons are the most common products, while aromatic hydrocarbons are generally not included. As used herein, the term "substantially inert" and "liquid medium" refers to a hydrocarbon or mixture of hydrocarbons which serves as a continuous phase in the reaction zone, as a means for suspending the catalyst particles, as a receptor for the heat of reaction, and as a carrier for the hydrocarbons produced. The liquid medium can react with the catalyst and reactive gases under the conditions of the reactor. This usually is not just a minor reaction, and generally most reactive gases form hydrocarbons without using a molecule of the liquid medium as a base. Both reactive gases and the product have some solubility in the substantially inert liquid medium. In effect, the product of the synthesis reaction is a mixture of hydrocarbons which may be distinguishable from and used as a liquid medium, or, alternatively, certain fractions of the product of the synthesis reaction may be used as a liquid medium. As used herein, the term "reaction zone" and "reaction volume" refer to a volume in the reactor comprising synthesis gas bubbles, solid particulate catalyst, and substantially inert liquids. Each of these components must have a fraction that is in turbulent motion. The invention is a process for producing hydrocarbons from hydrogen and carbon monoxide by reacting hydrogen and carbon monoxide in the presence of a particulate solid catalyst in a form that effectively controls the exothermic heat of the reaction. This reaction takes place in a container of the reactor adapted for the reaction of the gases in the presence of a substantially inert liquid medium and a bed of solid particulate catalyst. The hydrogen gas and the carbon monoxide gas, or more frequently a synthesis gas comprising both gases, are introduced into a plurality of places within the reactor vessel. Substantially inert liquids are also introduced into at least one place in the reactor and the flow of the liquid is substantially countercurrent to the gas flow. Bubbles of gas and substantially inert liquid medium flow upward through the zone containing solid catalyst particles at sufficient velocity to expand the zone to a volume greater than its static volume. The velocity of the fluids creates a single turbulent reaction where the liquid, the gas, and the solid catalyst are present and are in a state of motion. The velocity of the gas and the liquid are preferably controlled so that the reaction can proceed efficiently but the catalyst does not leave the zone with the flow or flows of gas and liquid. The reactor is a boiling bed reactor. The boiling bed comprises a liquid medium, catalyst particles and gas bubbles. The boiling bed has sufficient spatial velocity of the gas upstream of the gas bubbles and the liquid medium to cause the bed to expand to a volume greater than its static volume, thereby creating a turbulent reaction zone where the less something of the solid catalyst is in a state of motion. The gaseous space velocity, ie the gas yield divided by the total volume of the reactor containing the gas, should be between 200 and 20,000 m 3 per 3 catalyst per hour and preferably between 500 and 10,000 m 3 per m 3 of catalyst per hour. This spatial velocity allows a high conversion of synthesis gas to liquids. This spatial velocity of the gas is greater than that normally observed in boiling beds. This high performance and conversion is possible due to the heat removal efficiency of the preferred embodiments. The boiling beds also depend on the circulation of fluid to expand the bed. The liquid medium is introduced, flows substantially against the current with the gas, and this velocity. of the fluid increases the height of the boiling bed and the movement of the three phases. In the preferred embodiment the combined gas and liquid flows are sufficient to expand the bed to the point where a substantial portion, ie more than 90%, of the solid catalyst is in a state of motion, but the volume of the boiling bed is controlled to a predetermined amount of expansion.

The reactive gases can be introduced anywhere in the reactor as long as they come into contact with the catalyst and the liquid medium. Since the bubbles move upwards, it is usually advantageous to locate means for introducing the reactive gases near the bottom of a reactor to utilize the maximum volume of the reactor. The introduction of the gas near the bottom of the reactor can be by any means known in the art. Typically, means are selected such that the bubbles are formed with an average diameter of less than 5 millimeters, preferably less than 3 millimeters, and more preferably less than 1 millimeter. In addition to the distribution means, the bubble size is also a factor of the viscosity, surface tension and molecular weight of the liquid medium of the reactive fraction of the catalyst particles with the liquid medium. Those skilled in the art can easily influence the bubble size and vary the liquid properties, gas velocities, solid catalyst concentrations and gas distribution methods. Particularly preferred means for introducing gas into a transverse distributor plate introducing liquid medium and gas comprises the reactive gases uniformly through the bottom of the expanded bed to maintain a turbulent flow still stable within the bed. The transverse distributor plate also serves as a separation for separating the feed streams from substantially inert liquid and gas from the expanded catalyst bed and provides a physical support for the boiling catalyst bed within the reactor. These preferred means are described in U.S. Patent 4,874,583, which is incorporated herein by reference. Below this manifold plate, those skilled in the art can have sparger or sprinkler tubes or other mechanical devices to dissipate the kinetic energy of the fed gas and, if necessary, the liquid. This distributor plate injects fluids and gas into the reaction zone comprising the fluid, catalyst particles and gas bubbles. The reaction zone extends, without interruption, from about the lowest point of the reactor where the gas is injected into the bed containing liquid medium or solid particulate catalyst to approximately the substantially catalyst-free separation zone near the top of the reactor. reactor. Generally in large reactors the volume of the reactor is segregated by plates. Each plate can function to redistribute fluids and to support the catalyst. However, the substantial volume of the reactor is lost, because each plate generally must be preceded by a separation zone. In this invention there are no plates or other mechanical devices that can isolate the sections of the reactor but only at the expense of a substantial volume loss of the reactor. The entire volume within the uninterrupted reaction zone is available for catalyst particles in a liquid medium in contact with the gas bubbles. The weight of the catalyst load decreases as much by the buoyancy in the liquid medium as by the elevation created by the gas and liquid flowing upwards. Fluids are added from the bottom and extracted from the top of this reaction zone. The additional places where reactive gases are introduced are placed so that the gas is introduced into the existing reaction zone. The reaction zone does not include volume occupied by heat exchangers. The fraction of the reaction zone occupied by the solid catalyst ranges from about 10 to about 60 percent of the total reaction volume. A preferred embodiment has the fraction of the reaction zone occupied by solid catalyst in a range ranging from about 10 percent to about 50 percent of the total reaction volume. A more preferred embodiment has the fraction of the reaction zone occupied by solid catalyst in a range ranging from about 10 percent to about 40 percent of the total reaction volume. The substantially inert liquid medium and the gas bubbles will occupy the volume of the remaining reaction zone. The preferred fraction of the reaction zone occupied by the substantially inert liquid medium will be from about 30 percent to about 80 percent of the volume of the total reaction zone. The most preferred fraction of the reaction zone occupied by the substantially inert liquid medium will be from about 50 to about 70 percent of the volume of the total reaction zone. The fraction of the reaction zone normally occupied by the gases will be the rest not occupied by the solid catalyst or the liquid medium. The preferred fraction of the reaction zone occupied by the gas bubbles will be from about 5 to about 50 percent of the volume of the total reaction zone. The most preferred fraction of the reaction zone occupied by the gas will be from about 20 to about 45 percent of the volume of the total fire zone. The most preferred fraction of the reaction zone occupied by the gas will be from about 20 percent to about 40 percent of the volume of the reaction zone. Near the upper particle of the boiling bed reactor is a substantially catalyst-free zone which functions as a separation zone for the catalysis of the gaseous and liquid phases in the boiling bed. This separation zone is often about one tenth to about one third of the reactor volume. The liquid medium and the gases pass to the top to a hot separator where they can be treated outside the reactor. Depending on the amount of reactive gases present, one skilled in the art can recycle a portion of the separated gas through the reactor, or use this gas in a subsequent reactor, or use this gas as a fuel. Examples of treatment may include removing the product, removing some fraction from the liquid medium or by any means known in the art, cooling the medium, or filtering the fine particles from the medium. A preferred embodiment of this invention introduces this liquid medium treated back to the reactor. The reactor contains means for introducing gas comprising one or both reactive gases at a plurality of places within the single reaction zone. One such place is that the one discussed above near the bottom of the reaction zone, since the bubbles tend to rise through the bed. The other media can be located through the reaction zone. In a preferred embodiment, the synthesis gas injected into the bottom-most gas injection means has a molar ratio of hydrogen to carbon monoxide that is between 0.5: 1 to about 6.0: 1, preferably between 1.0: 1 to 3.0: 1; and more preferably between 1.6: 1 to 2.2: 1. This most preferred ratio generally goes a good reaction rate with a cobalt-based catalyst and gives greater selectivity to the desired heavier hydrocarbon product. The stoichiometric ratio of hydrogen to carbon monoxide used within the generation of waxy hydrocarbon may be in the range of about 2: 1 to about 3: 1 for the generation of methane. It is clear that when the reactants are converted, and some of the methane is inevitably generated, it will become deficient in a reactive gas, i.e., hydrogen, and the ratio of hydrogen to carbon monoxide moves out of the desired range in the absence of the supplemental gas . There are at least additional means for introducing gas to the reaction zone in a location above the lowest point where the gas is injected. Each of the additional gas introduction means is located within the reactor vessel at a predetermined distance from the bottom of the vessel. The exact locations of the addition means for introducing gas into the reaction zone will usually be chosen to maximize the efficiency of the reactor and the selectivity of the reaction. In a preferred embodiment, the additional means for introducing additional reactive gases into the single turbulent reaction zone is an apparatus comprising a sparger tube or sprayer attached to a tube. Bubbling or sprinkler tubes are designed and positioned to minimize back-mixing of reactive gases. Each of these means for introducing synthesis gas dissipates the gas at a plurality of points, and all those points are located at approximately the same distance from the bottom of the reactor vessel, they are approximately equidistant from a vertical axis defined by the centers of the horizontal cross sections of the reactor vessel, and are arranged with an approximate radial symmetry about the vertical axis. Proper placement of the sparger or spray tubes will reduce backmixing, encourage uniform distribution of gas bubbles, and control the growth of bubble size in the reaction zone. The reduction of backmixing in the boiling bed reactor will allow the unit to achieve an efficient conversion of hydrogen and carbon monoxide to product. The number of places where additional gas will be introduced will vary from about 1 to about 15. A preferred embodiment would contain from 2 to 9 places where additional gas is introduced into the reaction zone, while a more preferred mode would contain from 3 to 8 places, and the most preferred mode would contain 6 to 7 places. In the preferred embodiment the placement of those places and the amount of hydrogen gas or carbon monoxide gas or both introduced is sufficient to preserve a relatively constant ratio of hydrogen to carbon monoxide from about 0.5: 1 to about 6.0: 1, so more preferably from about 1.0: 1 to about 3.0: 1, and more preferably from about 1.6: 1 to about 2.2: 1, through the reaction zone. As the conversion of synthesis gas to hydrocarbons proceeds, the rate of reaction will fall as the hydrogen gas content of synthesis gas in the boiling bed is depleted. The carbon monoxide content of the synthesis gas in the boiling bed is also depleted, albeit at a lower speed than that of hydrogen. As the conversion proceeds, initial diluents such as carbon dioxide, nitrogen and methane become more prevalent. In addition, some of the F-T synthesis products, including lighter hydrocarbons and water, further dilute the reactive gases. The injection of refueling synthesis gas containing carbon monoxide gas and hydrogen can increase the concentration of both reagents in the synthesis gas and therefore increase the reaction rate. Suitable catalysts are those that promote a Fisher-Tropsh hydrocarbon synthesis. The particular catalyst selected depends on the preferred operating conditions and the desired product, and is typically selected, for example, from iron, cobalt, rare earth catalysts and mixtures thereof. A preferred catalyst for synthesizing waxy hydrocarbons is a cobalt and ruthenium catalyst. The active catalytic metal salts are impregnated or mixed with substrate particles such as alumina. The catalyst particles and their manufacture are known in the art. Such parass must be sufficiently large and sufficiently dense to have a tendency to settle in the substantially inert liquid. In the reaction zone this tendency will be counteracted by the movement of gas and fluid. In the separation zone, it is very desirable to cause the catalyst to sediment through the liquid and back into the reaction zone. For a boiling bed, it is important that the catalyst load does not contain too large a fraction of large particles or too large a fraction of small particles. An appropriate size distribution is important for the successful fluidization of the catalyst bed without entraining the catalyst with the extraction of liquid medium and product. One embodiment contains solid particulate catalyst having an average particle diameter of between about 0.2 and about 3.5 millimeters. A more preferred embodiment contains solid particulate catalyst having an average particle diameter of between about 0.3 and about 1.6 millimeters. A preferred size distribution will be narrow. Typically with a cobalt-based catalyst the contact between the synthesis gas and the catalyst results in the production of paraffinic hydrocarbons to a large extent, together with small amounts of olefins and oxygenates. The product, using such a cobalt-based catalyst, may contain about one-half by weight of hydrocarbons that boil in a moderate range, about one third by weight of waxy paraffin, and the rest of light boiling hydrocarbons and methane. The composition changes somewhat if iron-based catalysts are used. In any case, part of the product will evaporate towards the gas and will be transported to the reactor together with the inert compounds and the unconverted gas fed. The fraction of the product that is transported to the vapor phase will depend on the operating conditions that are being used. These hydrocarbons can be converted using methods known in the art. The remaining product will be miscible with the liquid medium. The substantially inert liquid medium comprises hydrocarbons with a boiling range above 340 ° C, ie paraffinic wax; that is to say, hydrocarbons with a boiling point range of about 200 ° C to about 340 ° C, here called mid-range boiling hydrocarbons; a product of the synthesis reaction, or mixtures thereof. The fraction of each can be varied to achieve an efficient operation. For exa, mid-boiling hydrocarbons minimize foam formation. The boiling point hydrocarbons in the middle range lower the solidification temperature of the mixture, and may allow the liquid to cool to a lower temperature than the waxy hydrocarbons. A preferred embodiment of the invention contains a substantially inert liquid comprised of from about 0.001% to about 50% mid-range hydrocarbons. Waxy paraffinic hydrocarbons can increase the viscosity of the liquid and are easy to separate from the lighter hydrocarbons. The properties of those fractions and the flow properties resulting from mixing them are known in the art. The liquid medium can also have a composition similar to that of the product. A portion of the reaction product will mix and may become distinguishable from the substantially inert liquid medium. Therefore, the product hydrocarbons will comprise a fraction of the liquid medium, and after a prolonged period the hydrocarbons in the product may comprise the majority of the liquid medium. For start-up purposes, if no product is available, poly-alpha-olefins or highly refined C30 to C50, ie hydrocarbons with heteroatoms and free aromatics, can be used as the substantially inert liquid medium. One embodiment of this invention requires the partial separation of those components from the extracted liquid medium, using separation methods known in the art, and reconstituting and recycling back to the reactor a fraction having the desired properties. Alternatively, an unfractionated fluid can be recycled to the boiling bed. A preferred embodiment will allow regulation of the temperature of the substantially inert liquid entering the reactor by eying an external heat exchanger to the reactor to cool the recirculating vapor. Typically, the liquid medium withdrawn from the reactor is obtained from the separation zone which is located above the reaction zone. Those skilled in the art will understand that it may also be advantageous to add, and extract fluid directly to, and from, the reaction zone. This will allow a greater control of the temperature and can provide a mechanism to remove or add catalyst to the bed.

The process pressure can be any pressure sufficient to allow the desired reaction. Typically, such a pressure is from about 1 atmosphere to about 100 atmospheres, preferably from 10 to 50 atmospheres, and more preferably from 20 to 50 atmospheres. Pressures less than 1 atmosphere require operation at vacuum conditions and which are not necessary and unduly expensive. Pressures greater than about 100 atmospheres would require significant costs due to the increased strength of the equipment necessary to withstand high pressures. The boiling bed is operated at a temperature sufficient for Fisher-Tropsh synthesis to occur. Although such temperatures depend on many factors, mainly the selected catalyst, a typical temperature in the reactor will be from about 150 ° C to 325 ° C, and more preferably from about 160 ° C to about 300 ° C, and so more preferably from about 160 ° C to about 240 ° C. The reaction rate at temperatures below 150 ° C would require an extraordinary amount of catalyst. Temperatures above about 325 ° C favor the generation of methane. The diameter of the boiling bed reactor is selected to give a surface velocity of gas fed (actual volumetric flow rate of the gases fed to reactor conditions divided by the empty cross-sectional area) between about 2 centimeters per second to about 30 centimeters per second. The most preferred range is from about 10 centimeters per second to about 20 centimeters per second. Gas velocity affects gas retention and gas-liquid mass transfer and back-mixing. The space velocity per hour of the preferred gas selected for optimum reactor efficiency should be between 200 and 20,000 m 3 per hour per m 3 of catalyst, more preferably between 500 and 10,000 m 3 per hour per m 3 of catalyst. For a boiling bed reactor the catalyst bed is partly supported by the flow of liquid medium. The rate of recycling of the liquid should be between about 0.1 and about 20 centimeters per second, more preferably between about 1 and about 10 centimeters per second. The space velocity per hour of the liquid should be between about 10 and about 100 volumes of liquid per volume of catalyst per hour. The space velocity per hour of the preferred liquid should be from about 20 to about liquid volumes per volume of catalyst per hour.

Due to the exothermic nature of the F-T reaction, the reactor is advantageously equipped with heat removal capabilities, so that the desired reaction temperature can be carefully controlled. In a boiling bed reactor, the heat can be removed in a variety of ways, including decreasing the sensible heat of the feed, internal heat exchangers, and external heat exchangers. The inclusion of each of these methods is a preferred embodiment of the invention. The preferred means for regulating the temperature of the reactor comprises employing an internal heat exchanger which comprises a tube. Those skilled in the art will know that heat exchangers are often used in multiple tubes, in series and in parallel, and those are included in this description. The tube may also have vanes or other devices known in the art to facilitate heat transfer, between the boiling bed and the cooling fluid. The cooling medium enters the tube from outside the reactor vessel. In a more preferred embodiment the cooling medium entering the tube comprises a liquid which is at least partially evaporated within the tube. Such cooling liquids include, for example, hydrocarbon-based liquids, commercially available Dowtherm, halogenated hydrocarbons and water. The most preferred embodiment uses Dowtherm, water, glycols, and mixtures thereof, which are introduced into the tube in liquid form and are evaporated at least partially before extracting those liquids from the tube. Other heat exchangers known in the art, including external heat exchange sleeves, may be included. Another preferred means for removing heat from the reactor is to introduce liquid medium to the reactor at a temperature lower than the temperature of the reactor. To the extent that the liquid medium introduced is the recycled liquid medium, a heat exchanger must be used to cool the liquid. Heat exchangers that remove heat from the liquid are well known in the art. In addition to the means described above for removing heat, the substantially inert liquid medium acts as a heat sink. With conventional gas phase processes, the heat released during the reaction substantially increases the temperature of the gas and the catalyst, which prevents the hydrocarbon synthesis reaction, increases the formation of methane, and causes deactivation or destruction of the catalyst. The high thermal capacity of the liquid phase allows high conversions while moderating the temperature gains.

The heat is transferred to the fluid and the fluid facilitates the removal of heat. The synthesis gases can be preheated before entering the reactor. Because the F-T reaction is highly exothermic, it is usually not necessary to heat the feed gas to the reaction temperature. The gas can also be heated by the substantially inert liquid in the reactor. The invention described above has many advantages. It provides several effective means to control the temperature, thus allowing greater conversions in the reactor without excessive methane formation or catalyst destruction. It provides means for controlling the ratio of hydrogen gas to carbon monoxide gas, thereby providing a mechanism for maximizing the reaction rates and the selectivity of the reaction through the reaction zone. It allows a high reaction bed height because the catalyst floats due to the gas and also possibly due to the liquid, which will reduce the cushioning of the catalyst for a given catalyst load. This invention also utilizes the maximum volume of the reactor, since plates that can be used to support catalyst and to redistribute fluids are not present in this invention. Another advantage is that a boiling bed has an almost zero pressure difference across the reaction zone, thereby providing a more uniform pressure that improves the productivity and selectivity to C5 + hydrocarbons. Figure 1 is a schematic diagram of a boiling bed reactor for the conversion of pressure from synthesis gas to liquid hydrocarbons. The reaction zone extends from this diffuser to approximately the point where the internal reactor heat exchangers exit to the catalyst bed. Above the reaction zone is a separation zone. An approximate place for the transition for the reaction zone to the separation zone is that shown by the arrow line opposite the wavy line 10. The synthesis gas enters the boiling bed 10 via line 22. The synthesis gas can to be preheated in the heat exchanger 24. The synthesis gas is introduced into the reaction zone via sparger tubes or sprinklers 16 located in or around the base plate 14 or the bottom of the reaction zone, or, alternatively, via a plate • transverse distributor also described as point 14. Three sparger tubes or single spacers are in place to add, in this embodiment, a gas comprising hydrogen gas to the reaction zone. The synthesis gas can be supplemented with hydrogen-enriched gas from line 20. Line 20 also supplies hydrogen enriched with gas to the reaction zone via sparger tubes or sprinklers 18 located in the reaction zone. A cooling coil or coils 12 can pass through the reaction zone, but without interfering with the introduction of the gas via the sparger or sprinkler tubes 16 and 18 or with the vertical migration of the fluid. Substantially inert fluid is introduced to the boiling bed reactor 10 via a boiling pump 56 and line 40. The fluid can flow up through the openings in the base plate 14 or alternatively the gas can be introduced via the plate transverse distributor 14. Liquid medium and gas leave the reactor and are separated from each other in the hot separator. Fluids and gas leave the boiling bed reactor 10 via line 42. If the separation zone is large then the fluids can leave the boiling bed as separate gas and fluid flows. In this embodiment, however, the fluids and gas then pass to a hot separator 44 which can be operated at a reduced pressure. By controlling the pressure and temperature of the hot separator, the composition of the recycled liquids can be varied. Substantially inert fluid leaves the hot separator via line 48, and a portion can be extracted as a product via line 50. The remaining fluid can be cooled with a heat exchanger 52 and then passed via line 54 back to the boiling pump 56. The catalyst can be added to the zone of reaction via, for example, line 30 and extracted from the reaction zone via line 32. A commercial modality of the FT synthesis process can be optimized to produce mainly fuels and lubricants as products. Figure 2 is a schematic of such a process, where the boiling bed reactor 1 is designed for a very high conversion to a single performance. The internal parts of the reactor are similar to those shown in Figure 1 and are not shown here. This embodiment of the invention shows the synthesis gas entering the reactor of the boiling bed 100 via line 140. Details of the boiling bed were omitted in this figure. The fluids leave the boiling bed reactor via line 102 and are optionally cooled in the heat exchanger 4. The fluids leave the heat exchanger via line 116 to the extraction vessel 118 and then via line 120 to the second one. separation vessel 122. The middle distillates are extracted via line 124, a portion is separated as product via line 126, and a portion is optionally recycled to the boiling bed via line 128 through heat exchanger 4, and through of the boiling pump 132. The liquids recycled via line 130 are a mixture, which can be varied, by controlling the relative flows of the liquid through lines 128 and 136. The liquids recycled via line 130 are preferably of a boiling point sufficiently greater than the reaction temperature to not produce a significant vapor phase. Another preferred commercial application of the F-T synthesis process mainly produces fuels as products. Figure 3 is a schematic of such a process where three reactors are used in series to process the fed synthesis gas. The F-T synthesis occurs in the first two reactors. This embodiment of the invention incorporates several reactors of the boiling bed. The synthesis gas enters the reactor of the boiling bed 202 via line 200. The details of the boiling bed were omitted in this figure. The fluids leave the boiling bed reactor via line 204 and then pass to the separation vessel 206. The light gases and naphthas are removed from the upper part of the separation vessel 206 via line 208. They are not shown in the figure. separation and removal of contaminants, mainly water vapor, between the vapors leaving the hot separator 4 via line 103. The water vapor content of more than 50% in the gas phase seriously prevents the FT synthesis process. The heavier hydrocarbons leaving the separation vessel 206 via the line 210 and a portion are transported to the heat exchanger 212. The fluids leave the heat exchanger to the boiling pump 214 and are recycled to the boiling bed 202. The gases Unconverted feeds are mixed with a fraction of liquid components and introduced into the second reactor of the boiling bed. A fraction of the heavier hydrocarbons of line 210 are optionally mixed with the gases and naphtha of line 208 and are fed via line 218 to the second reactor of boiling bed 220. The synthesis gas and the additional hydrogen enriched gas which feed the lines to reactor 220 are not shown. The fluids leave the boiling bed reactor 220 via line 222 and then pass through a heat exchanger 224. The cooled fluids pass via line 226 to separation vessel 228. Gases and light naphtha are extracted from the of the separation vessel 228 via line 230 and are separated as a product, used as fuel, or used for other processes. The heavier hydrocarbons leave the separation vessel 228 via line 232 and are further cooled in the heat exchanger 234. The fluids then pass via line 236 to a settling vessel 238 and then to a stripping vessel 240. The distillates means are removed from the top via line 242. A portion is separated as a product on line 244, and a portion passes via line 246 through heat exchanger 234 to be heated. This fluid when mixed with liquids from line 258 then passes through boiling pump 248 and is recycled via line 250 to boiling bed 220. The higher boiling fraction of the hot separator is mixed with hydrogen and steam and it is introduced to the boiling third reactor. The function of this third reactor is not that of an F-T synthesis but of converting the high molecular weight hydrocarbons into hydrocarbons of lower molecular weight. The heavy hydrocarbons leaving the separation vessel 240 via line 252 are supplemented with hydrogen and steam fed from line 252 and transported in catalytic thermofracting unit 260. The hydrocarbons leaving the catalytic thermoforming unit 260 via line 262 are fractionated in vessel 264 and gases and naphthas are extracted via line 266. The heavier hydrocarbons leave the separation vessel 262 via line 268 and a portion is extracted as middle distillates and a portion is recycled through the pump of boiling 264 to the catalytic thermofraction unit.

EXAMPLES As an example, 1000 standard cubic feet per hour of synthesis gas containing 46.8 mole percent nitrogen, 33.1 mole percent hydrogen, 16.8 mole percent carbon monoxide, 2.9 mole percent carbon dioxide, and 0.32 mole percent of water vapor are contacted with 0.61 cubic feet of a cobalt and ruthenium catalyst on an alumina support at an average reactor temperature of 204 ° C, a pressure of 370 psi (23649.72 kgf / m2) , and an hourly space velocity of 1,650 cubic feet per hour per cubic meter of catalyst. An 85% conversion of carbon monoxide to hydrocarbon under these conditions is achieved in a single pass through the reactor. The distribution of the resulting hydrocarbon product is shown below: Hydrocarbon Products lb / hr kg / hr% by weight Methane 0.45 0.204 8.08 Ethane 0.11 0.049 1.90 Propane 0.14 0.063 2.57 Hydrocarbon Products Normal Butane 0.17 0.77 3.04 Normal Pentane 0.19 0.086 3.42 C6-C12 1.51 0.6849 27.28 C13-C19 1.28 0.58 23.10 C20 + 1.70 0.77. 30.61 100

Claims (28)

CHAPTER CLAIMEDICATORÍO Having described the invention, it is considered as a novelty and, therefore, the content is claimed in the following CLAIMS:

1. In a process for producing hydrocarbons from hydrogen and carbon monoxide by reacting hydrogen and carbon monoxide in a reaction zone comprising solid particulate catalyst and substantially inert liquid medium, the improvement characterized in that it comprises: conducting the reaction in a turbulent reaction zone within a boiling bed reactor vessel adapted for the reaction of the gases in the presence of a substantially inert liquid medium and solid particulate catalyst, where the gas comprising one or more of the hydrogen gas and carbon monoxide gas is introduced into a plurality of places within the reaction zone, so that the bubbles of the gases flow upwardly through the reaction zone comprising solid catalyst particles and substantially inert liquid medium at a rate sufficient to expand the bed at a volume greater than its static volume.

2. The process according to claim 1, characterized in that the substantially inert liquid medium is added and extracted from the reaction zone, thereby creating a flow which facilitates the expansion of the bed to the point where at least some of the solid catalyst it is in a state of random movement.

3. The process according to claim 2, characterized in that each place where hydrogen gas or carbon monoxide gas is introduced, or both, is located inside the reactor vessel at a predetermined distance from the bottom of the container. .

The process according to claim 2, characterized in that the hydrogen gas, carbon monoxide gas, or mixture thereof is introduced into the plurality of predetermined distances from the bottom of the container.

5. The process according to claim 4, characterized in that at least one of the locations, where the location is determined by the distance from the bottom of the reactor, where hydrogen gas or carbon monoxide gas, or both gases are introduced and the gas is distributed to a plurality of points located at approximately the same distance from the reactor vessel, and approximately equidistant from the vertical axis defined by the center of a horizontal cross-section of the reactor vessel.

The process according to claim 1, characterized in that the gas is introduced through an apparatus which comprises a tubular portion extending inwards, through the wall of the reactor and a portion of sparger tube or sprayer, the sparger or sprayer portion is attached to the tubular portion and distributes gas to a location within the turbulent reaction zone.

7. The process according to claim 2, characterized in that the diameter of the bubbles of the synthesis gas is less than about 5 millimeters.

8. The process according to claim 2, characterized in that the liquid medium is selected from the group consisting of paraffin wax, hydrocarbons with a boiling point of about 150 ° C to about 340 ° C, and mixtures thereof.

9. The process according to claim 2, characterized in that there are means for modifying the liquid medium or catalyst, or both while the reaction is proceeding.

10. The process according to claim 2, characterized in that at least a portion of the liquid medium that is extracted is recycled.

The process according to claim 10, characterized in that the concentration of the average distilled fraction having a boiling point range of 200 ° C to about 340 ° C in the recycled liquid medium is from about 0.001 to about 50% in volume.

12. The process according to claim 1, characterized in that the solid particulate catalyst has an average particle diameter of from about 0.2 to about 3.5 millimeters.

The process according to claim 1, characterized in that the solid particulate catalyst has an average particle diameter of from about 0.3 to about 1.6 millimeters.

14. The process according to claim 4, characterized in that the molar ratio of hydrogen to carbon monoxide is in the reaction zone of about 0.5: 1 to about 6.0: 1.

15. The process according to claim 14, characterized in that the molar ratio of hydrogen to carbon monoxide is in the reaction zone of about 1.0: 1 to about 3.0: 1.

16. The process according to claim 15, characterized in that the molar ratio of hydrogen to carbon monoxide is in the reaction zone of about 1.6: 1 to about 2.2: 1.

17. The process according to claim 2, characterized in that the reactor vessel comprises means for removing heat.

18. The process according to claim 17, characterized in that the means for removing heat comprises a tube arranged so that cooling medium can enter and exit the tube out of the reactor vessel.

19. The process according to claim 18, characterized in that the cooling medium entering the tube comprises a liquid which evaporates at least partially inside the tube.

The process according to claim 19, characterized in that the cooling medium is selected from the group consisting of Dowthern, water, glycols, and mixtures thereof.

21. The process according to claim 10, characterized in that it also comprises means external to the reactor vessel for removing heat from the liquid medium that is recycled.

22. The process according to claim 2, characterized in that the space velocity of the gas is between about 200 and about 20,000 cubic meters per cubic meter per hour.

23. The process according to claim 22, characterized in that the space velocity of the gas is between about 500 and about 10,000 cubic meters per cubic meter catalyst per hour.

24. The process according to claim 2, characterized in that the liquid flow is between about 1 and about 10 centimeters per second.

25. The process according to claim 2, characterized in that the space velocity of the liquid is between about 10 and about 100 cubic meters of substantially inert liquid per cubic meter of catalyst per hour.

26. The process according to claim 2, characterized in that the space velocity of the liquid is between about 20 and about 80 cubic meters of substantially inert liquid per cubic meter of catalyst per hour.

27. The process according to claim 4, characterized in that the hydrogen gas, carbon monoxide gas, or a mixture thereof is introduced from 3 to 8 predetermined distance from the bottom of the container.

28. The process according to claim 4, characterized in that the hydrogen gas, carbon monoxide gas, or a mixture thereof is introduced from 6 to 7 predetermined distance from the bottom of the container.