Patent Application
20180029003
Not applicable.
This disclosure relates to systems and methods for producing Fischer-Tropsch hydrocarbons via a Fischer-Tropsch reactor cascade. More particularly, this disclosure relates to the production of Fischer-Tropsch hydrocarbons via a Fischer-Tropsch reactor cascade comprising at least one first Fischer-Tropsch reactor in fluid communication with at least one second Fischer-Tropsch reactor, wherein the at least one first Fischer-Tropsch reactor has a heat transfer surface area to catalyst volume ratio that is either less than or greater than that of the at least one second Fischer-Tropsch reactor.
The Fischer-Tropsch process was developed as a way of producing hydrocarbons from coal, natural gas, biomass, and the like. The formation of valuable products from natural gas (which may comprise primarily methane), coal, biomass, and other hydrocarbonaceous sources typically incorporates an intermediate step that entails the conversion of the hydrocarbonaceous source to synthesis gas or ‘syngas’, which is a mixture comprising carbon monoxide and hydrogen. The Fischer-Tropsch (or ‘FT’) process is a catalytic and exothermic process that is utilized to produce petroleum substitutes, typically gasoline-range boiling hydrocarbons used as automotive fuels. More recently, Fischer-Tropsch is increasingly being used as a method for preparing heavier hydrocarbons, such as diesel fuels, and waxy molecules that may subsequently be converted into desirable products, such as, but not limited to, lubricants.
The Fischer-Tropsch process involves a series of chemical reactions that produce a variety of hydrocarbons. FT reactions produce alkanes, for example, via the simplistically expressed Equation (1):
(2n+1)H2+nCO→CnH2n+2+nH2O, (1)
where ‘n’ is a positive integer. Thus, typical FT reaction products include paraffins, represented by the formula CnH2n+2, where n (i.e., the average carbon number of the product) is determined by reaction conditions including, but not limited to, temperature, pressure, space velocity, catalyst type, and feed stream composition. The formation of methane (i.e., n=1) is generally undesirable. A majority of the alkanes produced via the Fischer-Tropsch synthesis are straight-chain alkanes, although branched alkanes are also formed. In addition to alkane formation, competing reactions result in the formation of alkenes, as well as alcohols and other oxygenated hydrocarbons. In applications, relatively small quantities of non-alkane products are formed, although catalysts favoring some of these products have been developed.
A variety of catalysts can be used for the Fischer-Tropsch process, with the most common comprising the transition metals cobalt, and iron. Also useful are exotic metals like ruthenium. Nickel may be employed; however, nickel-based catalysts tend to favor the formation of methane, which is also known as ‘methanation’. Bimetallic Fischer-Tropsch catalysts, such as nickel-iron (e.g., Fe3Ni, FeNi3), cobalt-ruthenium, cobalt-platinum, cobalt-palladium, and the like, are also known in the art.
Cobalt-based catalysts are highly active, although iron-based Fischer-Tropsch catalysts are generally considered to be more suitable for low-hydrogen-content synthesis gases, such as those derived from coal, due to the tendency of iron-based catalysts to promote the water-gas-shift reaction (also referred to herein as the ‘WGSR’). In addition to the active metal or metals, Fischer-Tropsch catalysts typically contain a number of ‘promoters’, including, but not limited to, potassium and copper. Group 1 alkali metals, such as potassium, are typically considered to be a poison for cobalt catalysts, while being a promoter for iron catalysts. Fischer-Tropsch catalysts are often supported on high-surface-area binders/supports such as silica, alumina, titania, and zeolites.
As noted hereinabove, iron-based catalysts promote a water-gas-shift, which provides additional hydrogen via the reaction of Equation (2):
H2O+CO→H2+CO2 (2)
Accordingly, iron-based Fischer-Tropsch catalysts can generally tolerate feed streams comprising significantly lower molar ratios of hydrogen to carbon monoxide than can catalysts that do not promote or do not so highly promote the water-gas shift reaction (e.g., cobalt-based catalysts). This reactivity can be important for applications in which the synthesis gas for Fischer-Tropsch synthesis is derived from coal and/or biomass. Such synthesis gas tends to have relatively low molar ratios of hydrogen to carbon monoxide (e.g., less than or equal to about 1). Cobalt catalysts are typically more active for Fischer-Tropsch synthesis when the molar ratio of hydrogen to carbon monoxide in the feed synthesis gas is higher, such as when the feedstock synthesis gas is derived from natural gas. Synthesis gas produced from natural gas tends to comprise a higher molar ratio of hydrogen to carbon monoxide than the stoichiometric ratio of 2.1, so the water-gas-shift is typically not needed to enhance the molar ratio of such synthesis gas for use with cobalt-based catalysts. Iron-based catalysts are thus often preferred over cobalt-based catalysts for application with lower quality feedstocks, such as synthesis gas produced from coal and/or biomass.
Fischer-Tropsch catalysts deactivate by a variety of mechanisms. Catalyst deactivation and poisoning are caused by many factors, including, for example, undesired reaction of the active metal (e.g., reaction with sulfur). While, as a result of the water-gas-shift reaction, iron catalysts are generally preferred for use with lower quality feedstocks, these catalysts tend to form a number of undesirable products, including various oxides and carbides, and are well known to produce undesirably large amounts of carbon dioxide (e.g., via water-gas shift).
The utility of FT catalysts is decreased if they exhibit high methanation activity during FT synthesis. High levels of catalytic methane formation from carbon monoxide and hydrogen decreases the utility of a FT catalyst for formation of higher hydrocarbons. For example, the utility, as a Fischer-Tropsch catalyst, of nickel on conventional metal oxide supports is decreased as a result of the high methanation activity typical of nickel-based Fischer-Tropsch catalysts.
Cobalt-based catalysts are highly active, and, as mentioned hereinabove, are especially useful when the feedstock is formed from natural gas. Because of a high molar ratio of hydrogen to carbon monoxide typical of such feedstocks, water-gas-shift is not needed therewith. However, some feedstocks tend to also include sulfur-based components, and the sensitivity of the catalyst to sulfur may be significantly enhanced for cobalt-based catalysts relative to their iron counterparts, as cobalt-based Fischer-Tropsch catalysts often strongly adsorb sulfur. Furthermore, the cost of a cobalt-based catalyst may be more than ten times the cost of an iron-based catalyst. In extreme instances, virtually every atom of sulfur that enters the reactor may attach to a catalytically active site on a cobalt-based catalyst and poison it.
Techniques for removing sulfur from feedstock gas upstream of FT reactor(s) are known, and typically include the use of a vessel loaded with zinc oxide (or other suitable component/support). However, these systems require considerable external pressure loading and are expensive as a result of necessary compressor equipment, raw materials, and utilities.
Accordingly, there are needs in the art for enhanced systems and methods for the production of Fischer-Tropsch hydrocarbons. There are needs in the art for systems and methods to desirably provide for effective and economical reduction of the concentration of sulfur-based components and/or other impurities in a FT feed stream, whereby the lifetime of the FT catalyst can be extended. Desirably, such systems and methods provide for the affordable removal of sulfur and/or other impurities, thus enhancing the lifetime of the Fischer-Tropsch catalyst, and/or enhancing the productivity, activity, and/or selectivity thereof. In other situations, the synthesis gas used as a feed contains a very low concentration of sulfur compounds and other poisons and consequently there is less concern with catalyst poisoning and more concern with, and needs in the art to address, optimal catalyst utilization. There are also needs in the art for enhanced systems and methods to address situations wherein the synthesis feed gas has a high combined partial pressure of hydrogen and carbon monoxide.
There are disclosed herein one or more embodiments for a Fischer-Tropsch (“FT”) reactor system that includes a first FT reactor having a first FT catalyst and a first heat transfer surface area to catalyst volume ratio and being configured to receive a first feed comprising synthesis gas and, operating at first FT conditions, to convert a first portion of the synthesis gas in the first feed into first FT products, leaving unconverted a second portion of the synthesis gas. The first FT products comprise FT hydrocarbons. The FT reactor system includes a first separation apparatus configured to receive the first FT products as at least part of its feed and to separate the first FT products into first liquid FT hydrocarbons and a first FT tail gas stream comprising unreacted syngas. The FT reactor system further includes a second FT reactor, having a second FT catalyst and a second heat transfer surface area to catalyst volume ratio that is different from the first heat transfer surface area to catalyst volume ratio, in series with the first FT reactor. The second FT reactor is configured to receive a second feed comprising the first FT tail gas stream and, operating at second FT conditions, to convert at least a portion of the second feed into second FT products comprising second liquid FT hydrocarbons and a second FT tail gas stream.
The present disclosure also includes one or more embodiments of methods of producing FT hydrocarbons that includes the steps of introducing a first syngas feed comprising carbon monoxide and hydrogen into a first FT reactor having a first FT catalyst and a first heat transfer surface area to catalyst volume ratio, operating the first FT reactor at first FT operating conditions to convert a first portion of the syngas in the first syngas feed to FT product hydrocarbons, leaving a second portion of the syngas in the first syngas feed unconverted, separating the second portion of the syngas from liquid FT product hydrocarbons; introducing a second syngas feed comprising hydrogen and carbon monoxide and including the second portion of the syngas into a second FT reactor having a second FT catalyst and a second heat transfer surface area to catalyst volume ratio that is different from the first heat transfer surface area to catalyst volume ratio; operating the second FT reactor at second FT operating conditions to convert at least a portion of the syngas in the second feed to FT product hydrocarbons.
The present disclosure also includes one or more embodiments of methods of producing FT hydrocarbons that includes the steps of providing a carbonaceous source feed and converting the carbonaceous source feed to a first syngas feed, conditioning the first syngas feed into a first fresh syngas feed, forming at least a portion of a first FT feed, adjusting the temperature of the first FT feed, introducing the first FT feed into a first FT reactor stage comprising one or a plurality of FT reactors each having a first FT catalyst and a first heat transfer surface area to catalyst volume ratio; producing first FT hydrocarbon products in the first FT reactor stage operating under first FT operating conditions; separating the first FT hydrocarbon products into first liquid FT products and a first gas FT product stream; recycling a first portion of the first gas FT product stream as a portion of the first feed; using a second portion of the first gas FT product stream as at least part of a second FT feed; adjusting the temperature of the second FT feed; introducing the second FT feed having the adjusted temperature to a second FT reactor stage comprising one or a plurality of FT reactors each having a second FT catalyst and a second heat transfer surface area to catalyst volume ratio wherein a first ratio of the combined heat transfer surface area of all of the first FT reactors of the first FT reactor stage divided by the total combined catalyst volume of all of the first FT reactors of the first FT reactor stage differs from a second ratio of the combined heat transfer surface area of all of the second FT reactors of the second FT reactor stage divided by the total combined catalyst volume of all of the second FT reactors of the second FT reactor stage, operating the second FT reactor stage at second FT operating conditions to convert at least a portion of the syngas in the second feed to second FT product hydrocarbons, separating the second FT hydrocarbon products into second liquid FT products and a second gas FT product stream, recycling a first portion of the second gas FT product stream as part of the first FT feed, recycling a second portion of the second gas FT product stream as part of the second FT feed, adjusting the temperature of a third portion of the second gas FT product stream, separating the third portion of the temperature-adjusted second gas FT product stream into third liquid FT products and a third gas FT product stream, recycling a first portion of the third gas FT product stream as part of the first FT feed, recycling a second portion of the third gas FT product stream as part of the second FT feed, and recycling a third portion of the third gas FT product stream as part of the carbonaceous source feed.
The present disclosure also includes one or more embodiments of an apparatus comprising an FT reactor having a first FT catalyst and a first heat transfer surface area to catalyst volume ratio configured to receive a first feed comprising synthesis gas and to convert a first portion of the synthesis gas in the first feed into first FT products comprising FT hydrocarbons and leave unconverted a second portion of the synthesis gas. The FT reactor is further configured to provide the unconverted second portion of the synthesis gas to a second FT reactor having a second FT catalyst and a second heat transfer surface area to catalyst volume ratio that is different from the first heat transfer surface area to catalyst volume ratio.
The present disclosure also includes one or more embodiments of an apparatus comprising a Fischer-Tropsch (“FT”) reactor having a first FT zone configured to provide a first heat transfer surface area to catalyst volume ratio and a second FT zone configured to provide a second heat transfer surface area to catalyst volume that is different from the heat transfer surface area to catalyst volume ratio of the first zone, wherein the first FT zone has a first FT catalyst and is configured to receive a first feed comprising synthesis gas and to operate under first FT conditions to convert a first portion of the synthesis gas in the first feed into first FT products and leave unconverted a second portion of the synthesis gas and further configured to provide the unconverted second portion of the synthesis gas as at least a portion of a second feed to the second FT zone, and the second FT zone has a second FT catalyst and is configured to receive the second feed and to operate under second FT conditions to convert unconverted synthesis gas in the second feed into second FT products.
These and other embodiments, features and advantages will be apparent in the following detailed description and drawings.
For a more detailed description of the present invention, reference will now be made to the accompanying drawing, wherein:
As used herein, the term “tubular reactor” refers to Fischer-Tropsch reactors containing one or more tubes containing FT catalyst, wherein the inner diameter or average width of the one or more tubes is typically greater than about 0.5″.
As used herein, the phrase “a high-temperature Fischer-Tropsch (or ‘HTFT’) reactor” means an FT reactor that is typically operated at temperatures of 330° C.-350° C., which typically employs an iron-based catalyst. This process has been put to use extensively by Sasol in their Coal-to-Liquid (CTL) plants. As used herein, the phrase “a low-temperature Fischer-Tropsch (or ‘LTFT’) reactor” means an FT reactor that is operated at lower temperatures, generally in a range between 170° C.-235° C., which typically employs a cobalt-based catalyst.
As used herein, the term “microchannel reactor” refers to Fischer-Tropsch reactors containing one or more tubes or channels coated with and/or filled with Fischer-Tropsch catalyst, wherein the inner diameter or average width of the one or more tubes is less than about 0.25 inches, and more typically about 0.16 inches.
As used herein, the term “compact reactor” refers to Fischer-Tropsch reactors designed to operate at a higher specific cooling area. Such compact reactors generally have an inner diameter or average width of the one or more tubes that are larger than that of a microreactor but less than that of a conventionally sized FT reactor.
As used herein, the abbreviation “FT” and/or “F-T” stand for Fischer Tropsch (which may also be written “Fischer-Tropsch”).
As used herein, the term “FT tail gas” means gas produced from an FT reactor. The FT tail gas may typically contain unreacted hydrogen and carbon monoxide, as well as carbon dioxide, some light hydrocarbons, and other light reaction byproducts.
As used herein, the term “FT water” or “FT water stream” means water produced by an FT reaction. The FT water will typically include dissolved oxygenated species, such as alcohols, and light hydrocarbons.
As used herein, the term “liquid FT hydrocarbon products” means liquid hydrocarbons produced by an FT reactor.
As used herein, the term “poison” refers to a component that reversibly interacts with a particular catalyst (i.e., a catalyst inhibitor that slows the reaction rate) and/or irreversibly deactivates the catalyst. Examples include Group 1 alkali metals, such as potassium, and sulfur compounds with respect to cobalt-based catalysts and halides with respect to both cobalt-based and iron-based catalysts. As noted above, a substance may be a poison with respect to one type of catalyst while acting as a promoter for iron catalysts.
Use of the term “tubular” is not meant to be limiting to a specific cross sectional shape. For example, tubes may have a cross-sectional shape that is not circular. Accordingly, the tubes of a tubular reactor may, in one or more embodiments, have a circular, oval, rectangular, and/or other cross sectional shape(s).
As used herein, the terms “reformed gas,” “synthesis gas” and “syngas” are used to refer to streams comprising, but not limited to, hydrogen and carbon monoxide. When used to describe synthesis gas, the term “fresh” is used herein to indicate that the synthesis gas (i.e., the fresh synthesis gas) has not previously passed through an FT reactor and been extracted unreacted therefrom.
“Activity” is defined herein as a parameter that reflects the speed of conversion of carbon monoxide (CO) per unit catalyst and per unit time. This parameter may be expressed in such a way that is independent of temperature, pressure, and reactant concentration. It is usually expressed as the “A” in the equation:
wherein the rateCO is the rate of CO converted per unit volume of catalyst per unit time. EA is the Fischer Tropsch activation energy, R the gas Universal constant, and T the absolute temperature. The function of f(Pi) describes the impact of the reactant and product concentration on the rate of CO conversion and is expressed as a function of partial pressure (Pi). It may be difficult to compare two different catalysts because each may behave very differently under similar process conditions, depending on their composition and physical properties. Each catalyst composition will have its unique f(Pi) partial pressure function. The value of f(Pi) is usually obtained after much research working with several gas compositions and several temperature and pressure conditions.
“Productivity” is defined herein as the rate of carbon monoxide (CO) converted per unit time and per unit of catalyst at a predetermined temperature. Productivity may be expressed as standard volume of CO in cc converted per volume of catalyst in cc per hour.
“Selectivity” is defined herein as the fraction of a certain chemical compound produced from the overall carbon monoxide (CO) conversion.
The “catalyst volume” of a fixed bed in a reactor tube of a tubular FT reactor is defined as the total inner volume of that part of the reactor tube where the fixed bed of catalyst particles is present. The catalyst volume thus includes the total inner volume of that part of the reactor tube, both the volume occupied by the catalyst particles, as well as the volume comprised of the voids between the catalyst particles. The “total catalyst volume of a fixed bed, tubular reactor” means the sum of all of the catalyst volumes of all the tubes of that reactor.
As used herein and as mentioned above, the abbreviation “HTFT” with respect to an FT reactor stands for “high-temperature Fischer-Tropsch,” while the abbreviation “LTFT” with respect to an FT reactor stands for “low-temperature Fischer-Tropsch.”
As used herein, with respect to an FT plant, (1) the abbreviation “GTL” stands for gas-to-liquids; the abbreviation “CTL” stands for coal-to-liquids;
As used herein and as mentioned above, the abbreviation “WGSR” stands for water-gas-shift reaction.
As used herein, the abbreviation “S/V” stands for heat transfer surface area to catalyst volume ratio.
As used herein, the abbreviation “GHSV” stands for gas hourly space velocity.
As used herein, “carbonaceous” or “hydrocarbonaceous” feedstocks means hydrocarbon feedstocks used to make syngas and may include but are not limited to biomass, natural gas, associated gas, coal-bed methane, residual oil fraction(s), coal, brown coal, peat, municipal waste and combinations thereof.
Overview.
Herein disclosed are embodiments comprising systems, methods, and apparatuses for the production of Fischer-Tropsch products. One or more embodiments of this disclosure incorporates a cascade comprising at least two stages of Fischer-Tropsch reactors, wherein a first stage comprises at least one Fischer-Tropsch reactor having a first FT catalyst and a first heat transfer surface area to catalyst volume ratio, and a second stage comprises at least one Fischer-Tropsch reactor having a second FT catalyst and a second heat transfer surface area to catalyst volume ratio that Is different from the first heat transfer surface area to catalyst volume ratio. The first heat transfer surface area to catalyst volume ratio may be less than the second heat transfer surface area to catalyst volume ratio, which may be useful, for example, to protect the second FT catalyst from poisoning. Alternatively, the first heat transfer surface area to catalyst volume ratio may be greater than the second heat transfer surface area to catalyst volume ratio, which may be useful, for example, if the syngas feedstock being used for the first FT reactor is highly reactive, i.e, has a high combined partial pressure of hydrogen and carbon monoxide.
In one or more embodiments, there may be additional differences between the first and second FT reactors. For example, in one or more embodiments, the at least one Fischer-Tropsch reactor of the first stage is a different type of Fischer-Tropsch reactor than the at least one Fischer-Tropsch reactor of the second stage, e.g., the at least one Fischer-Tropsch reactor of the first stage may be a tubular reactor and the at least one Fischer-Tropsch reactor of the second stage may be a microchannel, compact reactor or slurry bed reactor. In one or more embodiments, the at least one Fischer-Tropsch reactor is a microchannel, compact reactor or slurry bed reactor and the at least one Fischer-Tropsch reactor of the second stage may be a tubular reactor. In one or more embodiments, the at least one Fischer-Tropsch reactor of the second stage has a substantially higher productivity than the at least one Fischer-Tropsch reactor of the first stage. In one or more embodiments, the at least one Fischer-Tropsch reactor of the second stage has a substantially lower productivity than the at least one Fischer-Tropsch reactor of the first stage. In one or more embodiments, the pressure drop over the at least one Fischer-Tropsch reactor of the second stage is greater than the pressure drop across the at least one Fischer-Tropsch reactor of the first stage. In one or more embodiments, the pressure drop over the at least one Fischer-Tropsch reactor of the second stage is less than the pressure drop across the at least one Fischer-Tropsch reactor of the first stage. These and other embodiments, and components of the systems and methods for producing FT products via the disclosed systems will be described in detail herein below.
Fischer-Tropsch System Comprising Reactor Cascade.
Herein disclosed are embodiments for a Fischer-Tropsch system comprising a cascade of at least two stages of FT reactors, each stage comprising at least one FT reactor. Description of the disclosed Fischer-Tropsch system will now be made with reference to
In addition to having different S/V ratios, the first FT reactor 100 may also differ from the second FT reactor 150 in other ways. For example, in one or more embodiments, the first FT reactor 100 is generally more resistant to poisoning by contaminants commonly found in a synthesis gas feed than is the second FT reactor 150. For example, the first FT reactor 100 may be more resistant to poisoning by sulfur compounds, including, but not limited to, hydrogen sulfide. For example, such poison resistance may be provided by have the first heat transfer surface area to catalyst volume ratio being smaller than the second heat transfer surface area to catalyst volume ratio or may be supplementally provided by using a poison-resistant catalyst in the first FT reactor (as the first FT catalyst). In this way, the first FT reactor 100 may dually serve as an FT production reactor and as a guard bed, protecting the second FT reactor 150 from poisoning. As opposed to conventional guard beds, however, the first FT reactor 100 produces FT products, i.e., the purpose of the first FT reactor 100 is not only to remove contaminants from a synthesis gas feed, but also to produce FT products. Due to the contaminant reduction provided by first FT reactor 100, the second FT reactor 150 may be operable with and/or may contain a more expensive catalyst than first FT reactor 100. The first FT reactor and the second FT reactor 150 may differ in other ways, for example, by productivity (which may include but is not limited to use of catalysts having a different level of productivity) and/or by operating temperature and/or by pressure drops, and/or by CO conversion levels and/or by water vapor partial pressures, as discussed further herein.
Still other examples of ways in which the first FT reactor 100 and the second FT reactor 150 may differ include by having a different pressure drop per unit reactor length and/or by use of a lower cost catalyst used in the first FT reactor 100 compared to the second FT reactor 150. In one or more embodiments, the first FT reactor 100 and the second FT reactor 150 may differ by dimension or their dimensions may be the same. In one or more embodiments, the first FT reactor 100 and the second FT reactor 150 may use different catalysts or their catalysts may be the same.
The First FT Reactor 100.
In embodiments, the first FT reactor 100 is an FT reactor of any type, having a first FT catalyst and a first heat transfer surface area to catalyst volume ratio that is different from a second heat transfer surface area to catalyst volume ratio of a second FT reactor in series with the first FT reactor. In one or more embodiments, the first FT reactor 100 comprises a fixed bed reactor. In one or more embodiments, the first FT reactor 100 comprises a tubular reactor. In one or more embodiments, the first FT reactor 100 comprises a fluidized bed reactor. In one or more embodiments, the first FT reactor 100 comprises a slurry bed reactor, such as, but not limited to, a slurry bubble column reactor. In one or more embodiments, the first FT reactor 100 comprises a microreactor or a compact reactor. In one or more embodiments, the first FT reactor 100 comprises an FT reactor of any type.
The disclosed Fischer-Tropsch reactor cascade system of
In one or more embodiments, the first FT reactor 100 contains and/or is configured for operation with a FT catalyst selected from cobalt-based FT catalysts, iron-based FT catalysts, or ruthenium-based FT catalysts. In one or more embodiments, the first FT reactor 100 contains and/or is configured for operation with a bimetallic FT catalyst. In one or more embodiments, the first FT reactor 100 contains and/or is configured for operation with a bimetallic FT catalyst selected from cobalt-ruthenium FT catalysts, iron-nickel catalysts, and combinations thereof. In one or more embodiments, the first FT reactor 100 contains and/or is configured for operation with a catalyst selected from cobalt-based catalysts and ruthenium-based catalysts. In one or more embodiments, the first FT reactor 100 does not contain and/or is not configured for operation with an Iron-based FT catalyst. In one or more embodiments, the first FT reactor 100 does not contain and/or is not configured for operation with an FT catalyst having iron as the predominant active metal. In one or more embodiments, the first FT reactor 100 contains and/or is operable with a catalyst that is less active (i.e., that produces less FT product per quantity of catalyst over time) than that of a second catalyst used with the second FT reactor 150. In one or more embodiments, the first FT reactor 100 contains and/or is operable with a low cost catalyst, i.e., a catalyst considered a sacrificial catalyst. Such an embodiment, however, is different from use of a guard bed, because the first catalyst used with the first FT reactor 100 would perform FT synthesis and the products produced by the first FT reactor 100 may be blended with products produced by the second FT reactor 150.
In one or more embodiments, the first FT reactor 100 contains and/or is operable with a catalyst exhibiting an activity towards production of heavy FT products. In such embodiments, the first FT reactor 100 has a higher selectivity of heavy products (carbon number 20 or higher) than the second FT reactor 150. This is a different consideration than methane selectivity. Preferentially, in such embodiments, the second FT reactor 150 has a selectivity of light products (C1-C20) of less than 70%. In one or more embodiments, the first FT reactor 100 contains and/or is operable with a low productivity cobalt-based FT catalyst. Suitable low productivity catalysts include, but are not limited to, Co/Si, Co/Ti, and Co/AI catalysts, with or without promoters. In one or more embodiments, the first FT reactor 100 is configured for operation at a productivity that is less than the productivity of the second FT reactor 150. In one or more embodiments, the first FT reactor 100 is configured for operation with a catalyst productivity of less than about 300, 250, or 200 standard cubic centimeters of carbon monoxide per cubic centimeter of catalyst per hour (cc CO/cc cat/h).
As previously mentioned herein, in one or more embodiments, the first FT reactor 100 comprises a tubular reactor or is tubular in nature. Accordingly, in one or more embodiments, the first FT reactor 100 comprises one or more reactor tubes having an FT catalyst disposed therein and/or thereon, as would be known to one of skill in the art. In one or more embodiments, the first FT reactor 100 comprises a tubular fixed bed FT reactor. The number of tubes in a multi-tubular reactor is not critical to the disclosure and may vary widely. In one or more embodiments, the first FT reactor 100 is a tubular reactor containing from about 0.15 to about 4 liters, from about 0.2 to about 3.5 liters, or from about 0.4 to about 3 liters catalyst coated and/or catalyst containing tubes. In one or more embodiments, the first FT reactor 100 comprises a tubular reactor containing tubes having an average tube length in the range from about 15 to about 40 feet, from about 15 to about 35, or from about 25 to about 30. In one or more embodiments, the first FT reactor 100 comprises a tubular reactor containing tubes having an average inner tube diameter (or average inner cross-section width) that is in the range from about 0.5 inch to about 2 inches, from about 0.6 to about 1.5 inches, or from about 0.7 to about 0.9 inch. In one or more embodiments, the first FT reactor 100 comprises a tubular reactor containing tubes having an average inner tube diameter (or average inner cross-section width) that is greater than or equal to about 0.5, 0.75, 1, or 2 inches. In one or more embodiments, the first FT reactor 100 comprises a tubular reactor containing at least one tube having an average inner tube diameter (or inner cross section width) that is greater than or equal to about 0.5, 0.75, 1, or 2 inches.
In one or more embodiments, the first FT reactor 100 has a heat transfer surface area to catalyst volume ratio that is less than the heat transfer surface area to catalyst volume ratio of the second FT reactor 150. In one or more embodiments, the configuration for the first FT reactor 100 having a lesser heat transfer surface area to catalyst volume ratio than the second FT reactor 150 enables the first FT reactor 100 to be operated with a lower CO conversion level, producing less liquid products than the second FT reactor 150, and to have a lower pressure drop than the pressure drop across the second FT reactor 150. In one or more embodiments, the first heat transfer surface area to catalyst volume ratio of the first FT reactor 100 is less than about 8 inch−1. In one or more embodiments, the first heat transfer surface area to catalyst volume ratio of the first FT reactor 100 is less than about 7.5 inch−1. In one or more embodiments, the heat transfer surface area to catalyst volume ratio of the first FT reactor 100 is less than about 7 inch−1.
In one or more embodiments, the first FT reactor 100 has a heat transfer surface area to catalyst volume ratio that is greater than the heat transfer surface area to catalyst volume ratio of the second FT reactor 150. Such embodiments may be useful, for example in situations where the first feed syngas for the first FT reactor has a high total partial pressure of hydrogen and carbon monoxide. For example, syngas produced from an autothermal reactor may typically have a higher total partial pressure of hydrogen and carbon monoxide than would syngas from a steam methane reformer. Use of such a syngas, having a higher total partial pressure of hydrogen and carbon monoxide, in an FT reactor results in generation of a large amount of heat, which is advantageously handled by an FT reactor having a high heat transfer surface area to catalyst volume ratio. The hydrogen and carbon monoxide in the unreacted synthesis gas exiting the first FT reactor 100 would have a lesser partial pressure than the hydrogen and carbon monoxide in the first feed syngas to the first FT reactor 100. The unreacted synthesis gas exiting the first FT reactor 100 would be introduced into the second FT reactor 150. The lower total partial pressure of hydrogen and carbon monoxide in the unreacted synthesis gas would be adequately handled by the lower heat transfer surface area to catalyst volume ratio of the second FT reactor 150. In one or more embodiments, the first FT reactor 100 is selected from non-tubular reactors and the second FT reactor 150 is a tubular reactor. In one or more embodiments, the first FT reactor 100 is selected from the group of microchannel reactors and compact reactors, while the second FT reactor 150 is a tubular fixed bed FT reactor. In one or more embodiments, the first FT reactor 100 comprises a slurry reactor, while the second FT reactor 150 is a tubular fixed bed FT reactor. Multi-tubular reactors suitable for the first FT reactor 100 in such applications include but are not limited to microchannel reactors described in U.S. Pat. No. 7,829,602, which is hereby incorporated herein by reference in its entirety for all purposes not contrary to this disclosure. In one or more embodiments, the first FT reactor 100 comprises a compact spiral plate and spiral tube reactor substantially similar to or the same as that described in U.S. Patent Application No. 61/799,485 or a compact spiral finned reactor substantially similar to or the same as that described in U.S. Patent Application No. 61/799,825, both of which were filed internationally as PCT/US14/29746 and each of which is incorporated herein by reference in its entirety for all purposes not contrary to this disclosure. In one or more embodiments, the first FT reactor 100 comprises a compact finned panel reactor substantially similar to or the same as that described in U.S. Patent Application No. 61/800,090, filed internationally as PCT/US14/29886, which is incorporated herein by reference in its entirety for all purposes not contrary to this disclosure.
In one or more embodiments wherein the first heat transfer surface area to catalyst volume ratio of the first FT reactor 100 is greater than the second heat transfer surface area to catalyst volume area of the second FT reactor 150, the first heat transfer surface area to catalyst volume ratio of the first FT reactor 100 is greater than about 8 inch−1. In one or more embodiments wherein the first heat transfer surface area to catalyst volume ratio is greater than the second heat transfer surface area to catalyst volume area, the first heat transfer surface area to catalyst volume ratio of the first FT reactor 100 is greater than about 8.5 inch−1. In one or more embodiments wherein the first heat transfer surface area to catalyst volume ratio is greater than the second heat transfer surface area to catalyst volume area, the first heat transfer surface area to catalyst volume ratio of the first FT reactor 100 is greater than about 9 inch−1. In one or more embodiments wherein the first heat transfer surface area to catalyst volume ratio is greater than the second heat transfer surface area to catalyst volume area, the first heat transfer surface area to catalyst volume ratio of the first FT reactor 100 is greater than about 8 inch−1, and the second heat transfer surface area to catalyst volume area of the second FT reactor 150 is less than about 8 inch−1. In one or more embodiments wherein the first heat transfer surface area to catalyst volume ratio is greater than the second heat transfer surface area to catalyst volume area, the first heat transfer surface area to catalyst volume ratio of the first FT reactor 100 is greater than about 8, 8.5, or 9 inch−, and the second heat transfer surface area to catalyst volume area of the second FT reactor 150 is less than about 8, 7.5, or 7 inch−1. In one or more embodiments wherein the first heat transfer surface area to catalyst volume ratio is greater than the second heat transfer surface area to catalyst volume area, the second heat transfer surface area to catalyst volume ratio of the second FT reactor 150 is less than about 8 inch−1. In one or more embodiments wherein the first heat transfer surface area to catalyst volume ratio is greater than the second heat transfer surface area to catalyst volume area, the second heat transfer surface area to catalyst volume ratio of the second FT reactor 150 is less than about 7.5 inch−1. In one or more embodiments wherein the first heat transfer surface area to catalyst volume ratio is greater than the second heat transfer surface area to catalyst volume area, the second heat transfer surface area to catalyst volume ratio of the second FT reactor 150 is less than about 7 inch−1.
In one or more embodiments, the first FT reactor 100 is operable at a lower gas hourly space velocity (GHSV) than the second FT reactor 150. In one or more embodiments, the first FT reactor 100 is configured for operation at a GHSV that is less than or equal to about 1000 h−1, less than or equal to about 1200 h−1, or less than or equal to about 1500 h−1.
In one or more embodiments, the first FT reactor 100 is configured for operation at a lower temperature than that for which the second FT reactor 150 is configured. In one or more embodiments, the first FT reactor 100 is configured for operation at a temperature in the range of from about 160° C. to about 240° C., from about 180° C. to about 235° C., or from about 190° C. to about 220° C. In one or more embodiments, the first FT reactor 100 is configured for operation at a pressure in the range of from about 200 psig to about 650 psig, from about 300 psig to about 480 psig, or from about 350 psig to about 450 psig.
In one or more embodiments, the first FT reactor 100 is configured for operation with a pressure drop thereacross that is less than the pressure drop for which the second FT reactor 150 is configured. In one or more embodiments, the first FT reactor 100 is configured for operation with a pressure drop of less than about 3 psi, less than about 2 psi, or less than about 1 psi per foot of reactor length. In one or more embodiments, the first FT reactor 100 is operable at a water vapor partial pressure that is less than that of the second FT reactor 150. In one or more embodiments, the first FT reactor 100 is operable at a water vapor partial pressure at the reactor exit of up to about 5, 4, or 3 bar.
The first FT reactor 100 produces a first FT water stream and first FT products comprising first liquid FT hydrocarbon products and a first FT tail gas stream. The first FT tail gas stream may include both gaseous FT hydrocarbon products, unreacted synthesis gas and in some cases other components. In one or more embodiments, the first FT water stream exits the first FT reactor 100 separately from the first liquid FT hydrocarbon products and the first FT tail gas stream. The feed to the second FT reactor comprises the at least a portion of the first FT tail gas stream.
The Second FT Reactor 150.
In embodiments, the second FT reactor 150 is an FT reactor of any type having a second FT catalyst and a second heat transfer surface area to catalyst volume ratio that is different from the first heat transfer surface area to catalyst volume ratio of the first FT reactor. In one or more embodiments, the second FT reactor 150 is a fixed bed reactor. In one or more embodiments, the second FT reactor 150 is a tubular reactor. In one or more embodiments, the second FT reactor 150 is a microchannel reactor or a compact reactor. In one or more embodiments, the second FT reactor 150 is a fluidized bed reactor. In one or more embodiments, the second FT reactor 150 is a slurry bed reactor, such as, but not limited to, a slurry bubble column reactor. In one or more embodiments, the second FT reactor 150 contains and/or is configured for operation with an FT catalyst selected from cobalt-based FT catalysts, iron-based FT catalysts, and ruthenium-based FT catalysts. In one or more embodiments, the second FT reactor 150 contains and/or is configured for operation with a bimetallic FT catalyst. In one or more embodiments, the second FT reactor 150 contains and/or is configured for operation with a bimetallic FT catalyst selected from cobalt-ruthenium FT catalysts, iron-nickel FT catalysts, cobalt-platinum FT catalysts, cobalt-palladium FT catalysts, and combinations thereof. In one or more embodiments, the second FT reactor 150 contains and/or is configured for operation with a catalyst selected from cobalt-based catalysts and ruthenium-based catalysts. In one or more embodiments, the second FT reactor 150 does not contain and/or is not configured for operation with an iron-based FT catalyst. In one or more embodiments, the second FT reactor 150 does not contain and/or is not configured for operation with an FT catalyst having iron as the predominant active metal. In one or more embodiments, the second FT reactor 150 contains and/or is operable with a higher productivity cobalt-based FT catalyst than is used with the first FT reactor 100. Suitable high productivity catalysts include, but are not limited to, Co/Ru, Co/Pd, and Co/Pt catalysts. As mentioned hereinabove, in one or more embodiments, the second FT reactor 150 is configured for operation at a catalyst productivity that is greater than the productivity of the first FT reactor 100. For example, the second FT reactor 150 may be configured for operation with a catalyst productivity of greater than about 300, 350, or 400 standard cubic centimeters of carbon monoxide per cubic centimeter of catalyst per hour (cc CO/cc cat/h). In one or more embodiments, the first FT reactor 100 is configured for operation with a catalyst productivity of less than about 300 cc CO/cc cat/h, while the second FT reactor 150 is configured for operation with a carbon monoxide conversion of greater than about 300 cc CO/cc cat/h. In one or more embodiments, the first FT reactor 100 is configured for operation with a catalyst productivity of less than about 300, 250, or 200 cc CO/cc cat/h, and the second FT reactor 150 is configured for operation with a catalyst productivity of greater than about 300, 400, or 600 cc CO/cc cat/h.
In one or more embodiments, the first FT reactor 100 is a tubular reactor and the second FT reactor 150 is selected from all other (non-tubular) types of FT reactors. In one or more embodiments, the second FT reactor 150 is selected from the group of microchannel reactors and compact reactors. In one or more embodiments, the second FT reactor 150 comprises a compact reactor. In one or more embodiments, the second FT reactor 150 comprises a tubular reactor or is tubular in nature. In one or more embodiments, the second FT reactor 150 comprises one or more reactor tubes and/or channels having FT catalyst disposed therein and/or thereon, as known to those of skill in the art. In one or more embodiments, the second FT reactor 150 is a tubular fixed bed FT reactor. Thus, in one or more embodiments, either or both of the first FT reactor 100 and/or the second FT reactor 150 are fixed bed tubular reactors. A reactor tube in either the first FT reactor 100 and/or the second FT reactor 150 may be filled partly or entirely with a catalyst bed comprising FT catalyst particles or filled with inert material of heat conductive material. As mentioned in the “Notation and Nomenclature” section above, the ‘catalyst volume’ of a fixed bed in a reactor tube is defined as the inner volume of that part of the reactor tube where the fixed bed of catalyst particles is present. This volume thus includes the both volume occupied by the catalyst particles, as well as the volume of the voids between the catalyst particles. In one or more embodiments, the first FT reactor 100 and/or the second FT reactor 150 comprise one or more reactor tubes with a fixed bed of catalyst particles over a predetermined length of the corresponding reactor tube.
As discussed in detail herein below, the second FT reactor 150 may comprise a multi-tubular reactor. To accommodate and provide for a second heat transfer surface area to catalyst volume ratio (also referred to herein as ‘S/V’), that is greater than the first heat transfer surface area to catalyst volume ratio of the first FT reactor 100, the second FT reactor 150 may, in one or more embodiments, be configured with a greater heat transfer surface area than that of the first FT reactor 100. This may result from, for example, the use of smaller but more numerous tubes within the second FT reactor 150 relative to the size and number of tubes in the first FT reactor 100. In one or more embodiments, the tubes in the second FT reactor 150 may be smaller in diameter than the tubes of the first FT reactor 100. As a result of the presence of a greater number of tubes in the second FT reactor 150 than in the first FT reactor 100, the spacing and dimensions of the tubes of the second FT reactor 150 may be smaller or compacted than tubes of the first FT reactor 100, such that the second catalyst is more compact than the first catalyst. The S/V ratio is inversely proportional to tube internal diameter, such that as the tube internal diameter decreases, the S/V ratio increases. This reduction in the tube's internal dimension may force to the use of smaller catalyst particle sizes and therefore may result in the greater pressure drop across the second FT reactor 150 relative to the pressure drop across the first FT reactor 100, as described further herein below.
Multi-tubular reactors and the use of same in Fischer-Tropsch systems and processes are known in the art. Multi-tubular reactors suitable for the second FT reactor 150 include but are not limited to microchannel reactors described in U.S. Pat. No. 7,829,602, which is hereby incorporated herein by reference in its entirety for all purposes not contrary to this disclosure. In one or more embodiments, the second FT reactor 150 comprises a microchannel reactor substantially similar to or the same as that described in U.S. Pat. No. 7,829,602, which is incorporated herein by reference in its entirety for all purposes not contrary to this disclosure. In one or more embodiments, the second FT reactor 150 comprises a compact spiral plate and spiral tube reactor substantially similar to or the same as that described in U.S. Patent Application No. 61/799,485 or a compact spiral finned reactor substantially similar to or the same as that described in U.S. Patent Application No. 61/799,825, both of which were filed internationally as PCT/US14/29746 and each of which is incorporated herein by reference in its entirety for all purposes not contrary to this disclosure. In one or more embodiments, the second FT reactor 150 comprises a compact finned panel reactor substantially similar to or the same as that described in U.S. Patent Application No. 61/800,090, filed internationally as PCT/US14/29886, which is incorporated herein by reference in its entirety for all purposes not contrary to this disclosure.
The number of channels in a microchannel reactor is not believed to be critical to the disclosure and may vary widely. In one or more embodiments, the second FT reactor 150 comprises a microchannel reactor containing channels having an average opening in the range from about 0.1 mm to about 8 mm, from about 0.2 mm to about 5 mm, or from about 0.5 mm to about 3 mm. In one or more embodiments, the second FT reactor 150 comprises a tubular or microchannel reactor containing at least one tube or microchannel having an average inner tube or microchannel opening that is less than about 0.5 Inches. In one or more embodiments, the first FT reactor 100 comprises a tubular reactor containing at least one tube having an average inner tube or diameter (or average cross sectional dimension) that is greater than about 0.5, 1, or 2 inches, while the second FT reactor 150 comprises a tubular or microchannel reactor containing at least one tube or microchannel having an average inner tube or microchannel opening (or average cross sectional dimension) that is less than about 3, 1, 0.2 mm.
As mentioned above, in one or more embodiments, the second FT reactor 150 has a second heat transfer surface area to catalyst volume ratio that is greater than the first heat transfer surface area to catalyst volume ratio of the first FT reactor 100. In one or more embodiments, the second heat transfer surface area to catalyst volume ratio of the second FT reactor 150 is greater than about 8 inch−1. In one or more embodiments, the second heat transfer surface area to catalyst volume ratio of the second FT reactor 150 is greater than about 8.5 inch−1. In one or more embodiments, the second heat transfer surface area to catalyst volume ratio of the second FT reactor 150 is greater than about 9 inch−1. In one or more embodiments, the first heat transfer surface area to catalyst volume ratio of the first FT reactor 100 is less than about 8 inch−1, and the second heat transfer surface area to catalyst volume area of the second FT reactor 150 is greater than about 8 inch−1. In one or more embodiments, the first heat transfer surface area to catalyst volume ratio is less than about 8, 7.5, or 7 inch−1, and the second heat transfer surface area to catalyst volume area is greater than about 8, 8.5, or 9 inch−1.
In alternative embodiments, the second heat transfer surface area to catalyst volume of the second FT reactor 150 may be less than the first heat transfer surface area to catalyst volume ratio of the first FT reactor 100. This may result from, for example, the use of larger but less numerous tubes within the second FT reactor 150 relative to the size and number of tubes in the first FT reactor 100. In one or more embodiments, the tubes in the second FT reactor 150 may be larger in diameter than the tubes of the first FT reactor 100. As a result of the presence of a lesser number of tubes in the second FT reactor 150 than in the first FT reactor 100, the spacing and dimensions of the tubes of the second FT reactor 150 may be lesser or less compacted than tubes of the first FT reactor 100, such that the first catalyst is more compact than the second catalyst. The S/V ratio is inversely proportional to tube internal diameter, such that as the tube internal diameter decreases, the S/V ratio increases. A reduction in the tube's internal dimension may force to the use of smaller catalyst particle sizes and therefore may result in the greater pressure drop across the first FT reactor 100 relative to the pressure drop across the second FT reactor 150.
In one or more embodiments, the second FT reactor 150 is operable at a higher gas hourly space velocity (GHSV) than the first FT reactor 100. In one or more embodiments, the second FT reactor 150 is configured for operation at a GHSV that is greater than or equal to about 1500 h−1, greater than or equal to about 2000 h−1, or greater than or equal to about 3000 h−1.
In one or more embodiments, the second FT reactor 150 is configured for operation at a higher temperature than that for which the first FT reactor 100 is configured. In one or more embodiments, the second FT reactor 150 is configured for operation at a temperature in the range of from about 200° C. to about 250° C., from about 190° C. to about 240° C., or from about 200° C. to about 230° C., while the first FT reactor 100 is configured for operation at a temperature in the range of from about 160° C. to about 240° C., from about 180° C. to about 235° C., or from about 190° C. to about 220° C. In one or more embodiments, the second FT reactor 150 is configured for operation at a pressure in the range of from about 200 psig to about 550 psig, from about 350 psig to about 500 psig, or from about 400 psig to about 450 psig.
In one or more embodiments, the second FT reactor 150 is configured for operation at a lower temperature than that for which the first FT reactor 100 is configured. In one or more embodiments wherein the second FT reactor 150 is configured for operation at a lower temperature than that for which the first FT reactor 100 is configured, the second FT catalyst is more active than the first FT catalyst.
In embodiments wherein a higher heat transfer surface area to catalyst volume is utilized in the second FT reactor 150 than the first FT reactor 100, in one or more embodiments, the second FT reactor 150 may be configured for operation with a smaller catalyst particle size resulting in a pressure drop thereacross that is greater than the pressure drop for which the first FT reactor 100 is configured. In one or more embodiments, the first FT reactor 100 is configured for operation with a pressure drop of less than about 3 psi, 2 psi, or 1 psi per foot of tube length and the second FT reactor 150 is configured for operation with a pressure drop of greater than about 4 psi, 8 psi, or 10 psi per foot of tube length.
Alternatively, in embodiments wherein a higher heat transfer surface area to catalyst volume is utilized in the first FT reactor 100 than in the second FT reactor 150, in one or more embodiments, the second FT reactor 150 may be configured for operation with a larger catalyst particle size resulting in a pressure drop thereacross that is less than the pressure drop for which the first FT reactor 100 is configured. In one or more embodiments, the second FT reactor 150 is configured for operation with a pressure drop of less than about 3 psi, 2 psi, or 1 psi per foot of tube length and the first FT reactor 100 is configured for operation with a pressure drop of greater than about 4 psi, 8 psi, or 10 psi per foot of tube length.
As mentioned herein above, in one or more embodiments, the second FT reactor 150 is operable at a water vapor partial pressure that is greater than that of the first FT reactor 100. In one or more embodiments, the second FT reactor 150 is operable at a water vapor partial pressure of up to at least about 6 bar. In one or more embodiments, the first FT reactor 100 is operable at water vapor partial pressures of less than about 5 bar, and the second FT reactor 150 is operable to water vapor partial pressures greater than 5 bar. In one or more embodiments, the first FT reactor 100 is operable at water vapor partial pressures of less than about 5 bar, and the second FT reactor 150 is operable to water vapor partial pressures of up to at least about 6 bar.
The second FT reactor 150 produces a second FT water stream and second FT products comprising second liquid FT hydrocarbon products and a second FT tall gas stream. The second FT tail gas stream comprises gaseous FT hydrocarbon products and unreacted synthesis gas and in some cases other components. In one or more embodiments, the second FT water stream exits the second FT reactor 150 separately from the second liquid FT hydrocarbon products and the second FT tail gas stream.
Separation Apparatus.
As depicted in
In one or more embodiments, the first separation apparatus 120 may comprise a gas/liquid separator. In one or more embodiments, in the first separation apparatus 120, liquid FT hydrocarbon products are separated from the first FT tail gas. The separated first liquid FT hydrocarbon products exit from the first separation apparatus 120 via a first separation apparatus liquid outlet line 125. The separated first FT tail gas stream exits the first separation apparatus 120 via a first separation apparatus gas outlet line 131.
The first separation apparatus 120 may comprise one or more gas/liquid separators, each of which may comprise any separator known in the art to be operable to separate liquid hydrocarbons from gaseous components within the FT product introduced thereto via line 115. For example, in one or more embodiments, the gas/liquid separator is selected from the group consisting of knock out drums, scrubbers or similar devices. The gas/liquid separator may remove hydrocarbon that can be condensed via cooling (as indicated by a cooler C1) and may reduce water content in the vapor stream. As the hydrocarbon is condensed, liquid drops are formed as a mist and are suspended in the vapor stream. The velocity of the vapor stream is reduced as the vapor stream enters the gas/liquid separator 120, causing the liquid drops to fall out of the vapor stream. As an alternative, contacting the vapor stream with a metal mesh or corrugated metal placed inside the gas/liquid separator 120 may force liquid drops onto a cold metal surface to enhance separation. In one or more embodiments, the first separation apparatus 120 may comprise a separator that washes the vapor stream, such as a gas/liquid contactor, a spray tower or a scrubber.
Similarly, a second separation apparatus 152 may in one or more embodiments comprise a gas/liquid separator, or a series of two or more gas/liquid separators, configured to separate the second liquid FT hydrocarbon products, extractable from the second separation apparatus 152 via the second FT product outlet line 117, from second FT tail gas stream extractable from the second separation apparatus 152 via a second separation apparatus gas outlet line 141. In one or more embodiments, in the second separation apparatus 152, the second liquid FT hydrocarbon products are separated from the second FT tail gases. The separated second liquid FT hydrocarbon products, exit from the second separation apparatus 152 via a second separation apparatus liquid outlet line 127. The separated second FT tail gas stream may exit the second separation apparatus 152 via the second separation apparatus gas outlet line 141. The second separation apparatus 152 may include a gas/liquid separator, which may comprise any separator known in the art to be operable to separate liquid hydrocarbons from gaseous components within the second FT product introduced thereto via the second FT product outlet line 117. The gas/liquid separator of the second separation apparatus 152 may be any separator known in the art to be operable to separate liquid hydrocarbons from gaseous components within the FT product introduced thereto via line 117. In one or more embodiments, the gas/liquid separator is selected from the group consisting of knock out drums and/or scrubbers.
Other System Components
As depicted in
In one or more embodiments, the synthesis gas production apparatus 40 is configured to produce syngas from a carbonaceous material selected from biomass, natural gas, associated gas, coal-bed methane, residual oil fraction(s), coal, and combinations thereof. In one or more embodiments, the syngas production apparatus 40 is configured to produce syngas from light hydrocarbons, including methane and/or other hydrocarbons in natural gas, by means of various reforming processes, including steam reforming, auto-thermal reforming, dry reforming, advanced gas heated reforming, and/or by partial oxidation (e.g., catalytic partial oxidation). In one or more embodiments, the syngas production apparatus 40 is configured to produce synthesis gas via the gasification of biomass and/or coal.
Continuing to refer to
In embodiments where the temperature adjuster H1 comprises a cooler, the cooler may operate via heat transfer from a cooled material (e.g., boiler feed water (“BFW”) entering on first temperature adjuster input line 1, as indicated in
In one or more embodiments, the temperature adjuster H1 comprises a cooler that lowers the temperature of the ‘dirty’ syngas introduced thereto via the syngas production apparatus outlet line 45 to a temperature of about 90° C., 80° C., or 70° C. or lower, for example, for a RECTISOL™ pre-wash.
In one or more embodiments, the temperature adjuster H1 comprises a heater that elevates the temperature of the ‘dirty’ syngas introduced thereto via the syngas production apparatus outlet line 45 to a temperature of at least about 200° C., 400° C., or 600° C., as is the case for the RTI-Eastman chemical technology.
Continuing to refer to
In one or more embodiments, the syngas clean-up apparatus 50 is configured to reduce the amount of hydrogen sulfide in the synthesis gas introduced thereto. In one or more embodiments, the syngas clean-up apparatus 50 is configured to reduce the amount of carbon dioxide in the synthesis gas introduced thereto. The syngas clean-up apparatus 50 may comprise, for example, an add gas removal unit operable to reduce the level of hydrogen sulfide, ammonia, and/or carbon dioxide in the synthesis gas introduced thereto. In one or more embodiments, syngas clean-up apparatus 50 comprises one or more apparatus selected from the group consisting of zinc oxide beds, SELEXOL® units, and RECTISOL™ units.
In one or more embodiments, the syngas clean-up apparatus 50 comprises a zinc oxide bed configured for the removal of hydrogen sulfide via adsorption thereof. In one or more embodiments, the syngas clean-up apparatus 50 comprises a SELEXOL® unit. SELEXOL® units operate via a physical separation that does not rely on a chemical reaction. The SELEXOL® solvent is an add gas removal solvent (specifically, a mixture of dimethyl ethers of polyethylene glycol) frequently utilized to separate acid gases such as hydrogen sulfide and carbon dioxide from feed gas streams (i.e., syngas), such as, but not limited to, those produced via the gasification of coal, coke, and/or heavy hydrocarbon oils. In one or more embodiments, the syngas clean-up apparatus 50 comprises a RECTISOL™ unit. Like SELEXOL® units, RECTISOL™ units operate via a physical, rather than a chemical, separation. RECTISOL™ units utilize methanol as a solvent to separate acid gases such as hydrogen sulfide and carbon dioxide from valuable feed gas streams. RECTISOL™ units are frequently utilized to treat gas streams (i.e., syngas) produced by the gasification of coal and/or heavy hydrocarbons, as the methanol solvent is operable to remove trace contaminants such as ammonia, mercury, and hydrogen cyanide commonly present in such product gas streams. In one or more embodiments, the syngas clean-up apparatus 50 is configured to reduce the level of H2S in the synthesis gas extracted therefrom, for example via a fresh synthesis gas feed line 105, to less than about 1, 0.5 or 0.1 ppm. In one or more embodiments, the syngas clean-up apparatus 50 is configured to reduce the level of CO2 in the synthesis gas extracted therefrom, for example via the first fresh synthesis gas feed line 105, to less than about 5000, 1000 or 500 ppm. The synthesis gas exiting from the syngas clean-up apparatus 50 may be referred to as “clean” or “fresh” synthesis gas.
In one or more embodiments, the syngas clean-up apparatus 50 is includes equipment to adjust the amount of hydrogen in the synthesis gas. For example, the syngas clean-up apparatus 50 may comprise a membrane designed for removing excess hydrogen from the synthesis gas introduced thereto.
In one or more embodiments, the fresh synthesis gas feed line 105 is configured to introduce a first feed syngas, including at least a first portion of the fresh synthesis gas to the first FT reactor 100, either directly, or as indicated in
As mentioned above, at least the first portion of the fresh synthesis gas in the fresh synthesis gas feed line 105 is introduced into the first FT reactor 100. A second fresh synthesis gas line 107 may be configured to introduce a second portion of fresh synthesis gas from the fresh synthesis gas feed line 105 into the second FT reactor 150. In this manner, and as discussed further herein below, the molar ratio of hydrogen to carbon monoxide in the feed streams to the first FT reactor 100 and/or the second FT reactor 150 can be maintained at a desired value, e.g., just below stoichiometric.
As described above, in operation, the first FT reactor 100 has a first FT catalyst and a first S/V that differs from a second S/V of a downstream second FT reactor having a second catalyst. In embodiments, the first and second catalysts may be different of the same.
Although the embodiment(s) depicted in
Continuing to refer to
In one or more embodiments, the Fischer-Tropsch reactor cascade system of the present disclosure may further comprise a second recycle line 143 fluidly connecting a second separation apparatus gas outlet line 141 with either or (as depicted in
As discussed further herein below, in embodiments each of the following may be adjusted, singly or in one or more combinations, to provide a desired molar ratio of hydrogen to carbon monoxide in the feed to each FT reactor: (1) the extent (if any) of recycle (a) to the first FT reactor 100 via lines 135 and/or 145, and/or (b) to the second FT reactor 150 via line 144; (2) the amount of fresh synthesis gas (a) introduced to the first FT reactor 100 via the fresh synthesis gas feed line 105 and/or (b) introduced into the second FT reactor 150 via the synthesis gas fresh feed line 107, and/or (3) the amount of hydrogen or other gas (if any) (a) Introduced into the first FT reactor 100 via the first supplemental gas line 36 and/or (b) introduced into the second FT reactor 150 via the second supplemental gas line 37.
For embodiments wherein the second heat transfer surface area to catalyst volume of the second FT reactor 150 is greater than the first heat transfer surface area to catalyst volume ratio of the first FT reactor 100, the introduction of fresh synthesis gas directly to the second FT reactor 150 increases the risk of poisoning the second catalyst used for the second FT reactor 150 since the fresh synthesis gas has not gone through the first FT reactor 100. In such embodiments, a conservative approach may be warranted with respect to how much fresh synthesis gas is introduced into the second FT reactor 150.
The amount of fresh synthesis gas introduced into the first FT reactor 100 via the fresh synthesis gas feed line 105 and the amount (if any) of fresh synthesis gas introduced into the second FT reactor 150 via the second fresh synthesis gas line 107 may be adjusted as desired. For example, for embodiments wherein the second heat transfer surface area to catalyst volume of the second FT reactor 150 is greater than the first heat transfer surface area to catalyst volume ratio of the first FT reactor 100, should the synthesis gas contain relatively little contaminant, a greater portion of the fresh synthesis gas may be directed to the second FT reactor 150 than when the fresh synthesis gas comprises substantial contaminants.
In one or more embodiments wherein the first S/V is less than the second S/V, the first FT reactor 100 is operated to provide from about 0 to about 100 percent, from about 10 to about 90 percent, from about 10 to about 80 percent, or from about 10 to about 70 percent of the overall carbon monoxide converted. In one or more embodiments, the second FT reactor 150 is operated to provide from about 0 to about 100 percent, from about 10 to about 90 percent, from about 20 to about 90 percent, or from about 30 to about 90 percent of the overall carbon monoxide converted. In one or more embodiments, the first FT reactor 100 is operated to provide less than or equal to about 50 percent of the overall carbon monoxide converted and the second FT reactor 150 is operated to provide the other more than or equal to about 50 percent of the overall carbon monoxide converted. In one or more embodiments, the first FT reactor 100 is operated to provide less than or equal to about 40 percent of the overall carbon monoxide converted and the second FT reactor 150 is operated to provide the other more than or equal to about 60 percent of the overall carbon monoxide converted. In one or more embodiments, the first FT reactor 100 is operated to provide less than or equal to about 30 percent of the overall carbon monoxide converted and the second FT reactor 150 is operated to provide the other more than or equal to about 70 percent of the overall carbon monoxide converted. Depending on the make-up of available synthesis gas, the first FT reactor 100 or the second FT reactor 150 may be utilized to provide the entirety of the conversion at various times. The amount of conversion effected by each FT reactor may be altered depending on the available synthesis gas feed.
Although not depicted in the embodiment(s) of
In one or more embodiments, the LFTL separator 122 may be configured to separate a FT tail gas from condensed LFTL (“CLFTL”) (and in such embodiments may be referred to as a “CLFTL separator 122”). The CLFTL extracted via the LFTL product line 167 (which may be called in such embodiments, the “CLFTL product line” 167) may be characterized by a hydrocarbon mixture obtain when using a refrigerant like propane, which allows condensation of C3, C4, C5, C6 linear and isomers, especially when the condensation takes place under pressure. In one or more embodiments, the LFTL separator 122 is replaced by an oil contactor so that the CLFTL is dissolved in re-circulated oil. The FT tail gas extracted via the LFTL separator output line 166 may be characterized by unreacted hydrogen, carbon monoxide, methane, carbon trioxide and light hydrocarbons. The hydrocarbon contact depends on the temperature of the chiller C3 and the subsequent LFTL separator 122. Typically, the FT tail gas extracted via the LFTL separator output line 166 will contain methane, ethane, ethylene, propane and propene, with small amounts of butanes and pentanes. In one or more embodiments, the CLFTL's are condensed at a temperature of less than or equal to about 4° C., 0° C., or −10° C. In such a case, a water removal step may be necessary to avoid water freezing within the equipment. Within a CLFTL separator 122, CLFTL's are separated from an uncondensed FT tail gas. As mentioned hereinabove, it is envisioned that a portion of the FT tail gas extracted from the CLFTL separator 122 via the LFTL separator output line 166 may be recycled to the first FT reactor 100, the second FT reactor 150, synthesis gas production apparatus 40, or a combination thereof. CLFTL may be extracted from CLFTL separator 122 via CLFTL product line 167.
As noted hereinabove, FT products produced via the disclosed system and method may be further upgraded as known in the art. For example, in one or more embodiments, the method further comprises upgrading and/or separating one or more desired products produced in the first FT reactor 100 and/or the second FT reactor 150. In embodiments, hydrotreatment and distillation are utilized to provide desired products from the FT products in the first separation apparatus liquid outlet line 125, the second separation apparatus liquid outlet line 127, the CLFTL product outlet line 167, or a combination or any two or all three thereof. In one or more embodiments, the FT liquid product produced in the first FT reactor 100 and/or the second FT reactor 150 is upgraded to provide one or more products selected from primarily FT naphtha, primarily FT diesel, FT drilling fuel, primarily FT jet fuel, primarily lubricants, or a combination of any two or more of FT naphtha, FT diesel, FT jet fuel, lubrication oils and FT wax.
In one or more embodiments and as mentioned above, from the syngas clean-up apparatus 50, the first feed gas, which might include additions from the first supplemental gas line 36, the first recycle line 135, and/or the fourth recycle line 145, passes through the second temperature adjuster H2, which may be configured for heating the first feed gas prior to being introduced into the first FT reactor 100. The first FT reactor 100 operates at suitable FT conditions with a first FT catalyst to produce first FT products from the feed gas including clean synthesis gas. To achieve suitable FT conditions, the first FT reactor 100 is operable with a heat transfer apparatus (as previously discussed with respect to the first steam drum 101 of
The first FT products produced in the first FT reactor 100 comprise liquid hydrocarbons, vaporous hydrocarbons and unreacted synthesis gas. The first FT products of the first FT reactor 100 comprises a substantial quantity of high molecular weight hydrocarbons, generally from about C5 to about C100, or larger. The liquid FT products are a mixture of hydrocarbons that is the result of a block of —CH2— and it grows with a growth probability factor called an alpha value between 0.8 to 0.97, according to the Anderson-Schulz-Flory distribution.
As depicted in
As described above, within the separation apparatus 120, the liquid FT hydrocarbon products are separated from first FT tail gas. In one or more embodiments, as depicted in
The first liquid FT hydrocarbon products are extracted from the first separation apparatus 120 via the first FT liquid product line 125. In one or more embodiments, the first liquid FT hydrocarbon products extracted via the first FT liquid product line 125 comprise primarily C5+ hydrocarbons. In one or more embodiments, with more than one set of the cooler-separator combination, the first separation apparatus 120 will condense a heavier cut of the first liquid FT hydrocarbon products than the subsequent separators in series with the first separation apparatus 125.
The first FT tail gas separated from the first liquid FT hydrocarbon products within the first separation apparatus 120 is extracted in a stream via a first separation apparatus gas outlet line 131. The first FT tail gas stream extracted from the first separation apparatus 120 via the first separation apparatus gas outlet line 131 comprises unreacted synthesis gas and may further comprise carbon dioxide, and/or low molecular weight hydrocarbons. In one or more embodiments, the unreacted synthesis gas of the first FT tail gas extracted via the first separation apparatus gas outlet line 131 has a molar ratio of hydrogen to carbon monoxide that is in the range of from about 0.9 to about 2.2, from about 1.2 to about 2, or from about 13 to about 1.7. In one or more embodiments, the unreacted synthesis gas of the first FT tail gas extracted via the first separation apparatus gas outlet line 131 has a molar ratio of hydrogen to carbon monoxide that is greater than or equal to about 0.7:1, greater than or equal to about 0.8:1, greater than or equal to about 0.9:1, or greater than or equal to about 1:1. In one or more embodiments, the unreacted synthesis gas of the first FT tail gas extracted via the first separation apparatus gas outlet line 131 comprises less than about 10 ppb, 2 ppb, or 0.1 ppb hydrogen sulfide, as the first FT reactor 100 has served not only to as a production reactor, producing FT hydrocarbons, but has also served to clean the unreacted synthesis gas prior to introduction into the second FT reactor 150.
In one or more embodiments, a first portion of the separated first FT tail gas stream in the first separation apparatus gas outlet line 131 may be recycled to become part of the feed for the first FT reactor 100 via the first recycle line 135. At least a second portion of the separated first FT tail gas stream from the first separation apparatus 120 is introduced via the first separation apparatus gas outlet line 131 into the second FT reactor 150, as a least a part of a second feed gas. In addition to the second portion of the separated first FT tail gas stream, the second feed gas may include the second portion of the fresh synthesis gas feed in the second fresh synthesis gas line 107, at least a portion of a second FT tail gas stream recycled from the second FT reactor 150 after separation by a second separation apparatus 152 via a second recycle line 143 and a third recycle line 144, additional gas (such as, but not limited to, FT tail gas from the LFTL separator output line e 166, and/or hydrogen and/or nitrogen) in a second supplemental feed line 37, or a combination thereof, which may be combined with the synthesis gas in line 131. In one or more embodiments, hydrogen and/or nitrogen may be introduced into the second FT reactor 150 via the second supplemental feed line 37. In one or more embodiments, no fresh synthesis gas is introduced into the second FT reactor 150 via the second fresh synthesis gas line 107, and all of the fresh synthesis gas is passed through the first FT reactor 100.
In one or more embodiments and as mentioned above, from the first separation apparatus 120, the second portion of the fresh synthesis gas, with or without additions from the second supplemental feed line 37, the second fresh synthesis gas line 107, and/or the third recycle line 144, passes through a third heater H3 configured for heating the second feed gas prior to the second feed gas being introduced into the second FT reactor 150.
As depicted in of
As with the first FT reactor 100, the second FT reactor 150 is operable with the second FT catalyst at second FT conditions with a second heat transfer apparatus configured to maintain a desired reaction temperature, as known in the art. For example, as indicated in the embodiment of
In one or more embodiments, the second FT reactor 150 operates at suitable second FT conditions with a second FT catalyst (the first FT reactor being operated with the first FT catalyst) to produce second FT products from the second feed. As previously mentioned, the second FT products produced in the second FT reactor 150 comprise second liquid FT hydrocarbons, a second FT tail gas stream and a second FT water stream. The second FT tail gas stream typically comprises vaporous FT hydrocarbons and unreacted synthesis gas. In one or more embodiments, the second FT hydrocarbons extracted from the second FT reactor 150 via the second FT product outlet line 117 are cooled by a second cooler C2 prior to separation of liquids and gases within the second separation apparatus 152. For example, the second FT hydrocarbons may be cooled from a FT reaction temperature of greater than or equal to about 180° C., 200° C., or 220° C. to a temperature of less than or about equal to 100° C., 25° C., or 10° C. The second liquid FT products are extracted from the second separation apparatus 152 via a second FT liquid product line 127. In one or more embodiments, the second liquid FT products extracted via the second FT liquid product line 127 comprise primarily C5 hydrocarbons. The second cooler C2 may operate via the transfer of heat from the second liquid FT hydrocarbon products and the second FT tail gas to a coolant (e.g., BFW). The coolant may be introduced to the second cooler C2 via a second BFW line 9 and may be discharged via a second cooler discharge line 10.
In one or more embodiments as depicted in
As noted hereinabove, the fifth recycle line 133 may fluidly connect first separation apparatus 120 with separator 122, whereby a portion of the first gas FT products separated from the first liquid FT products in the first separation apparatus 120 may be introduced into the separator 122. The second separation apparatus gas outlet line 141 and the second feedline 142 may fluidly connect the second separation apparatus 152 with the separator 122, whereby a portion of the second gas FT products separated from the second liquid FT products in the second separation apparatus 152 may be introduced into the separator 122. As mentioned hereinabove, the output of the separator 122 may be fluidly connected with the syngas production apparatus 40, whereby at least a portion of the FT tall gas extracted from the separator 122 via the LFTL separator output line 166 may be utilized by the syngas production apparatus 40 to produce additional synthesis gas. Alternatively or additionally, the separator 122 may be fluidly connected with inputs of the first FT reactor 100 and/or the second FT reactor 150, whereby at least a portion of the FT tail gas extracted via the LFTL separator output line 166 may be introduced into the first FT reactor 100, the second FT reactor 150, or both, for example, via the first supplemental gas line 36 and/or the second supplemental gas line 37.
One or more supplemental gas lines, such as but not limited to first and/or second supplemental gas lines 36, 37, may be utilized to introduce additional feed gas into the first FT reactor 100 and/or the second FT reactor 150, respectively. For example, the first supplemental gas line 36 and/or the second supplemental gas line 37 may be configured to introduce hydrogen, nitrogen, FT tail gas (e.g., via the LFTL separator output line 166) into the first FT reactor 100 and/or the second FT reactor 150, respectively. The introduction of additional hydrogen via the first supplemental gas line 36 and/or the second supplemental gas line 37 may be utilized to provide a desired molar ratio of hydrogen to carbon monoxide in the feed gas introduced into the first FT reactor 100 and/or the second FT reactor 150. In one or more embodiments, the Fischer-Tropsch reactor cascade system of the present disclosure may further comprise a hydrogen supply apparatus (not depicted in
The Fischer-Tropsch reactor cascade system of the present disclosure may further comprise downstream FT product upgrading equipment. As such product upgrading equipment is known in the art, details of such equipment will not be provided here. For example, in one or more embodiments, the Fischer-Tropsch reactor cascade system of the present disclosure further comprises one or more apparatus selected from hydrolsomerisers, hydrotreaters, distillation apparatus, and other known upgrading and/or separation apparatus, configured to upgrade and/or separate FT products into one or more desirable components, including, but not limited to, FT diesel, FT jet, FT naphtha and FT waxes. For example, upgrading apparatus may be configured to produce paraffins from olefins in one or more of the FT product streams within the first separation apparatus liquid outlet line 125, the second separation apparatus liquid outlet line 127, and the LFTL product line 167.
The Fischer-Tropsch reactor cascade system of the present disclosure may further comprise various heaters and coolers, as known in the art. For example, as mentioned hereinabove, the Fischer-Tropsch reactor cascade system of the present disclosure may comprise a temperature adjuster H1, which may comprise a heater exchanger configured to heat or cool the ‘dirty’ synthesis gas prior to introduction into syngas clean-up apparatus 50. The Fischer-Tropsch reactor cascade system of the present disclosure may also comprise the second temperature adjuster H2 configured to heat the feed gas introduced into the first FT reactor 100, and/or a third temperature adjuster H3 configured to heat the feed gas introduced into the second FT reactor 150, or a combination thereof. As noted hereinabove, the Fischer-Tropsch reactor cascade system of the present disclosure may comprise a first cooler C1 configured to cool the FT products extracted from the first FT reactor 100 prior to separation in the first separation apparatus 120, a second cooler C2 configured to cool the FT products extracted from the second FT reactor 150 prior to separation in the second separation apparatus 152, a third cooler C3 configured to cool the synthesis gas containing stream(s) introduced thereto via the second feedline 142 and/or the fifth recycle line 133 prior to introduction thereof into the separator 122, or a combination of two or more of coolers C1, C2, and C3.
As described hereinabove, the first and second FT reactors 100 and 150 comprise heat transfer apparatus for maintaining a desirable operation temperature therein, as well known in the art. For example, the first FT reactor 100, the second FT reactor 150, or both may contain heat transfer tubes configured for the introduction of coolant thereto, whereby reaction heat is transferred to the coolant in the tubes and extracted from the reactor(s). For example, such heat transfer tubes may be configured to introduce water or some other suitable fluid into the reactor and for the extraction of steam therefrom. Such a heat transfer system may be associated with a reactor steam drum 101, 102, as known in the art, and described hereinabove. As such heat transfer apparatus is well known in the art, same will not be described herein in detail. In an alternative configuration the catalyst could be placed inside the tubes and the coolant be flowing in the outer shell of the reactor. The water vaporization would take place then in the shell side and transfer to the steam drum.
Other equipment, known in the art and not depicted in
As discussed herein above, although the description has been made with respect to a single the first FT reactor 100 having a first FT catalyst and a first S/V upstream of a single downstream the second FT reactor 150 having a second FT catalyst and a second S/V that is different than the first S/V, it is to be understood that either the first and/or the second stage of the disclosed system may comprise a plurality of reactors. The plurality of reactors in each stage may be arranged in series and/or in parallel. For example, in one or more embodiments, a first stage may comprise two or more first FT reactors 100, each of the plurality of the first FT reactors 100 of the first stage being substantially as described with regard to the first FT reactor 100 hereinabove. In one or more embodiments, a first stage comprises a plurality of first FT reactors 100 aligned in series. In one or more embodiments, a first stage comprises a plurality of first FT reactors 100 arranged in parallel. In this manner, it may be feasible to have one or more first FT reactor(s) 100 of the plurality of first FT reactors 100 in the first stage online while another or more of the plurality of first FT reactors 100 is being subjected to catalyst rejuvenation, regeneration, or replacement. Similarly, a second stage of FT reactors 150 may be utilized, each of the plurality of second FT reactors 150 being substantially similar to the second FT reactor 150 described hereinabove. In one or more embodiments, a second stage of FT reactors comprises a plurality of second FT reactors 150 aligned in series. In one or more embodiments, a second stage of FT reactors comprises a plurality of second FT reactors 150 arranged in parallel. In this manner, it may be feasible to have one or more second FT reactor(s) 150 of the plurality of second FT reactors 150 in the second stage online while another or more of the plurality of second FT reactors 150 is being subjected to catalyst rejuvenation, regeneration, or replacement. In one or more embodiments, a first stage of first FT reactors comprises more FT reactors than a second stage of second FT reactors, with the synthesis gas separated from the first FT product of the first FT reactors of the first stage being routed into the smaller number of second FT reactors of the second stage.
It is also envisaged that, in one or more embodiments, the Fischer-Tropsch reactor cascade system of the present disclosure may comprise, instead of two (or more) FT reactors, a single FT reactor configured with multiple zones. For example, in one or more embodiments, a single, multi-zoned FT reactor comprises a first zone configured to provide a heat transfer surface area to catalyst volume ratio as described with regard to the first FT reactor 100 hereinabove and a second zone configured to provide a heat transfer surface area to catalyst volume as described with regard to the second FT reactor 150 hereinabove that is different from the heat transfer surface area to catalyst volume ratio of the first zone. In one or more embodiments, a single multi-zoned FT reactor comprises a first zone configured to provide a productivity (cc CO converted/cc catalyst/hour) as described with regard to the first FT reactor 100 hereinabove and a second zone configured to provide a productivity (cc CO converted/cc catalyst/hour) as described with regard to the second FT reactor 150 hereinabove.
The locations of introduced and withdrawn streams indicated in
Methods for Producing FT Product Via FT Reactor Cascade.
Also disclosed herein are methods of producing FT hydrocarbons via a FT reactor cascade. Referring now to
For example, in one or more embodiments, the method of
Referring again to
Continuing to refer to
In one or more embodiments, the second synthesis gas feed is comprised solely of the first gas FT products. In one or more embodiments, the second synthesis gas feed is comprised of the first gas FT products and of other inputs for example, as described with respect to
In the one or more methods of
Turning now to
Accordingly, the first syngas feed may be conditioned 302 into a first fresh syngas feed, which forms at least a part of a first FT feed. For example, the “dirty” synthesis gas from the syngas production apparatus may be introduced into a syngas clean-up apparatus. The dirty syngas may be heated or cooled using a temperature adjuster depending on the type of cleanup system used in the syngas clean-up apparatus. In one or more embodiments, the syngas clean-up apparatus may comprise a first unit to perform a wash step. In such embodiments, the temperature of the dirty syngas would need to be cooled, if the syngas comes to the syngas clean-up apparatus directly from the syngas production apparatus. In one or more embodiments, the first unit of the syngas clean-up apparatus is an acid gas removal unit that operates below room temperatures. In such embodiments, the temperature of the dirty syngas would need to be cooled to a temperature close to the wash stream.
In one or more embodiments, where the carbonaceous source feed includes desulfurized natural gas and the dirty syngas comes from a steam methane reformer, then the dirty syngas may need to be cooled prior to contacting the hydrogen membrane. In one or more other embodiments, the dirty syngas may need to be heated prior to contacting an adsorbent bed, such as a zinc oxide bed. The ‘clean’ synthesis gas extracted from syngas clean-up apparatus may contain less than or about 10, 5, or 1 ppb volume percent hydrogen sulfide. Although indicated as a single apparatus, it is to be understood that syngas clean-up apparatus may comprise more than one unit.
The conditioned first syngas feed forms at least a portion of a first FT feed. The temperature of the first FT feed may be adjusted 307, if needed. The first gas feed is introduced 310 into a first FT reactor having a first FT catalyst and a first heat transfer surface area to catalyst volume ratio. The first FT reactor, operating under FT conditions, produces 320 first FT hydrocarbon products from the first feed. The first FT reactor is operable with a heat transfer apparatus (which may be such as previously discussed with respect to the first steam drum 101 of
The first FT hydrocarbon products are separated 330 into first liquid FT products and first gas FT products. The first liquid FT products are sent 332 for further processing, and/or to storage and/or to be transported offsite. A first portion of the first gas FT products may be recycled 335, forming a portion of the first FT feed.
Referring to
Continuing to refer to
Although recycle of various synthesis gas and FT tail gas stream(s) is described herein, in one or more embodiments, the system and method may be operated as a once-through system and/or method in certain applications. In other applications, recycle of one or more synthesis gas streams (e.g., recycle of synthesis gas from the first separation apparatus 120 to the first FT reactor 100 via the first recycle line 135, recycle of synthesis gas from the second separation apparatus 152 to the first FT reactor 100 via the fourth recycle line 145, recycle of synthesis gas from the second separation apparatus 152 to the second FT reactor 150 via the third recycle line 144) and/or recycle of one or more tail gas streams (e.g., recycle of FT tail gas from the LFTL (or CLFTL) separator 122 to the first FT reactor 100, the second FT reactor 150, and/or the syngas production apparatus 40 via the LFTL separator output line 166) is employed to enhance the overall production of liquid hydrocarbons via the disclosed system and method.
While preferred embodiments of the invention have been shown and described, modifications thereof can be made by one skilled in the art without departing from the spirit and teachings of the invention. The embodiments described herein are exemplary only, and are not intended to be limiting. Many variations and modifications of the invention disclosed herein are possible and are within the scope of the invention. Where numerical ranges or limitations are expressly stated, such express ranges or limitations should be understood to include iterative ranges or limitations of like magnitude falling within the expressly stated ranges or limitations. The use of the term ‘optionally’ with respect to any element of a claim is intended to mean that the subject element is required, or alternatively, is not required. Both alternatives are intended to be within the scope of the claim. Use of broader terms such as comprises, includes, having, etc. should be understood to provide support for narrower terms such as consisting of, consisting essentially of, comprised substantially of, and the like.
Accordingly, the scope of protection is not limited by the description set out above but is only limited by the claims that follow, that scope including all equivalents of the subject matter of the claims. Each and every claim is incorporated into the specification as an embodiment of the present invention. Thus, the claims are a further description and are an addition to the preferred embodiments of the present invention. The inclusion or discussion of a reference is not an admission that it is prior art to the present invention, especially any reference that may have a publication date after the priority date of this application. The disclosures of all patents, patent applications, and publications cited herein are hereby incorporated by reference, to the extent they provide background knowledge; or exemplary, procedural or other details supplementary to those set forth herein.
| Filing Document | Filing Date | Country | Kind |
|---|---|---|---|
| PCT/US16/19402 | 2/24/2016 | WO | 00 |
| Number | Date | Country | |
|---|---|---|---|
| 62120746 | Feb 2015 | US |
