Patent Application
20170081599
1. The Field of the Invention
The invention relates to heavy oil hydroprocessing methods and systems, such as ebullated bed hydroprocessing methods and systems, which utilize a dual catalyst system and operate at increased reactor severity.
2. The Relevant Technology
There is an ever-increasing demand to more efficiently utilize low quality heavy oil feedstocks and extract fuel values therefrom. Low quality feedstocks are characterized as including relatively high quantities of hydrocarbons that nominally boil at or above 524° C. (975° F.). They also contain relatively high concentrations of sulfur, nitrogen and/or metals. High boiling fractions derived from these low quality feedstocks typically have a high molecular weight (often indicated by higher density and viscosity) and/or low hydrogen/carbon ratio, which is related to the presence of high concentrations of undesirable components, including asphaltenes and carbon residue. Asphaltenes and carbon residue are difficult to process and commonly cause fouling of conventional catalysts and hydroprocessing equipment because they contribute to the formation of coke. Furthermore, carbon residue places limitations on downstream processing of high boiling fractions, such as when they are used as feeds for coking processes.
Lower quality heavy oil feedstocks which contain higher concentrations of asphaltenes, carbon residue, sulfur, nitrogen, and metals include heavy crude, oil sands bitumen, and residuum left over from conventional refinery process. Residuum (or “resid”) can refer to atmospheric tower bottoms and vacuum tower bottoms. Atmospheric tower bottoms can have a boiling point of at least 343° C. (650° F.) although it is understood that the cut point can vary among refineries and be as high as 380° C. (716° F.). Vacuum tower bottoms (also known as “resid pitch” or “vacuum residue”) can have a boiling point of at least 524° C. (975° F.), although it is understood that the cut point can vary among refineries and be as high as 538° C. (1000° F.) or even 565° C. (1050° F.).
By way of comparison, Alberta light crude contains about 9% by volume vacuum residue, while Lloydminster heavy oil contains about 41% by volume vacuum residue, Cold Lake bitumen contains about 50% by volume vacuum residue, and Athabasca bitumen contains about 51% by volume vacuum residue. As a further comparison, a relatively light oil such as Dansk Blend from the North Sea region only contains about 15% vacuum residue, while a lower-quality European oil such as Ural contains more than 30% vacuum residue, and an oil such as Arab Medium is even higher, with about 40% vacuum residue. These examples highlight the importance of being able to convert vacuum residues when lower-quality crude oils are used.
Converting heavy oil into useful end products involves extensive processing, such as reducing the boiling point of the heavy oil, increasing the hydrogen-to-carbon ratio, and removing impurities such as metals, sulfur, nitrogen and coke precursors. Examples of hydrocracking processes using conventional heterogeneous catalysts to upgrade atmospheric tower bottoms include fixed-bed hydroprocessing, ebullated-bed hydroprocessing, and moving-bed hydroprocessing. Noncatalytic upgrading processes for upgrading vacuum tower bottoms include thermal cracking, such as delayed coking, flexicoking, visbreaking, and solvent extraction.
Disclosed herein are methods for upgrading an ebullated bed hydroprocessing system to increase the rate of production of converted products from heavy oil. Also disclosed are upgraded ebullated bed hydroprocessing systems formed by the disclosed methods. The disclosed methods and systems involve the use of a dual catalyst system comprised of a solid supported catalyst and well-dispersed (e.g., homogeneous) catalyst particles. The dual catalyst system permits an ebullated bed reactor to operate at higher severity compared to the same reactor using only the solid supported catalyst.
In some embodiments, a method of upgrading an ebullated bed hydroprocessing system to increase rate of production of converted products from heavy oil, comprises: (1) operating an ebullated bed reactor using a heterogeneous catalyst to hydroprocess heavy oil at initial conditions, including (i) an initial reactor severity and (ii) an initial rate of production of converted products; (2) thereafter upgrading the ebullated bed reactor to operate using a dual catalyst system comprised of dispersed metal sulfide catalyst particles and heterogeneous catalyst; and (3) operating the upgraded ebullated bed reactor at (iii) a higher reactor severity and (iv) an increased rate of production of converted products than when initially operating the ebullated bed reactor.
In some embodiments, operating at higher severity includes: increasing throughput of heavy oil and operating temperature of the ebullated bed reactor while maintaining or increasing conversion of the heavy oil than when operating the ebullated bed reactor at the initial conditions. In other embodiments, operating at higher severity includes increasing conversion of heavy oil and operating temperature of the ebullated bed reactor while maintaining or increasing throughput of the heavy oil than when operating the ebullated bed reactor at the initial conditions. In yet other embodiments, operating at higher severity includes increasing conversion, throughput of heavy oil, and operating temperature of the ebullated bed reactor than when operating the ebullated bed reactor at the initial conditions.
In some embodiments, an increased throughput of heavy oil is at least 2.5%, 5%, 10%, or 20% higher than when operating the ebullated bed reactor at the initial conditions. In some embodiments, the increased conversion of heavy oil is at least 2.5%, 5%, 7.5%, 10%, or 15% higher than when operating the ebullated bed reactor at the initial conditions. In some embodiments, the increased temperature is at least 2.5° C., 5° C., 7.5° C., or 10° C. higher than when operating at the initial conditions. It will be appreciated, however, that in specific cases the exact temperature increase required to achieve the desired increase in rate of production of converted products can depend on the type of feedstock being processed and may vary somewhat from the temperature levels listed above. This is due to differences in the intrinsic reactivity of different types of feedstocks.
In some embodiments, the dispersed metal sulfide catalyst particles are less than 1 μm in size, or less than about 500 nm in size, or less than about 250 nm in size or less than about 100 nm in size, or less than about 50 nm in size, or less than about 25 nm in size, or less than about 10 nm in size, or less than about 5 nm in size.
In some embodiments, the dispersed metal sulfide catalyst particles are formed in situ within the heavy oil from a catalyst precursor. By way of example and not limitation, the dispersed metal sulfide catalyst particles can be formed by blending a catalyst precursor into an entirety of the heavy oil prior to thermal decomposition of the catalyst precursor and formation of active metal sulfide catalyst particles. By way of further example, methods may include mixing a catalyst precursor with a diluent hydrocarbon to form a diluted precursor mixture, blending the diluted precursor mixture with the heavy oil to form conditioned heavy oil, and heating the conditioned heavy oil to decompose the catalyst precursor and form the dispersed metal sulfide catalyst particles in situ.
In some embodiments, a method of upgrading an ebullated bed hydroprocessing system to increase rate of production of converted products from heavy oil comprises: (1) operating an ebullated bed reactor using a heterogeneous catalyst to hydroprocess heavy oil at initial conditions, including (i) an initial throughput, (ii) operating temperature, (iv) initial rate of production of converted products, and (iv) initial rate of fouling and/or sediment production; (2) thereafter upgrading the ebullated bed reactor to operate using a dual catalyst system comprised of dispersed metal sulfide catalyst particles and heterogeneous catalyst; and (3) operating the upgraded ebullated bed reactor at a higher throughput, a higher operating temperature, an increased the rate of production of converted products, and at a rate of fouling and/or sediment production equal to or less than when operating at the initial conditions.
In some embodiments, a method of upgrading an ebullated bed hydroprocessing system to increase rate of production of converted products from heavy oil comprises: (1) operating an ebullated bed reactor using a heterogeneous catalyst to hydroprocess heavy oil at initial conditions, including (i) an initial conversion, (ii) an initial operating temperature, (iii) an initial rate of production of converted products, and (iv) an initial rate of fouling and/or sediment production; (2) thereafter upgrading the ebullated bed reactor to operate using a dual catalyst system comprised of dispersed metal sulfide catalyst particles and heterogeneous catalyst; and (3) operating the upgraded ebullated bed reactor to hydroprocess heavy oil at a higher conversion, a higher operating temperature, an increased rate of production of converted products, and at a rate of fouling and/or sediment production equal to or less than when operating at the initial conditions.
These and other advantages and features of the present invention will become more fully apparent from the following description and appended claims, or may be learned by the practice of the invention as set forth hereinafter.
To further clarify the above and other advantages and features of the present invention, a more particular description of the invention will be rendered by reference to specific embodiments thereof which are illustrated in the appended drawings. It is appreciated that these drawings depict only typical embodiments of the invention and are therefore not to be considered limiting of its scope. The invention will be described and explained with additional specificity and detail through the use of the accompanying drawings, in which:
The present invention relates to methods for upgrading an ebullated bed hydroprocessing system to increase the rate of production of converted products from heavy oil and upgraded ebullated bed hydroprocessing systems formed by the disclosed methods. The methods and systems include (1) using a dual catalyst system and (2) operating an ebullated bed reactor at higher reactor severity to increase the rate of production of converted products.
By way of example, a method of upgrading an ebullated bed hydroprocessing system to increase rate of production of converted products from heavy oil, comprises: (1) operating an ebullated bed reactor using a heterogeneous catalyst to hydroprocess heavy oil at initial conditions, including (i) an initial reactor severity and (ii) an initial rate of production of converted products; (2) thereafter upgrading the ebullated bed reactor to operate using a dual catalyst system comprised of dispersed metal sulfide catalyst particles and heterogeneous catalyst; and (3) operating the upgraded ebullated bed reactor at (iii) a higher reactor severity and (iv) an increased rate of production of converted products than when initially operating the ebullated bed reactor.
The term “heavy oil feedstock” shall refer to heavy crude, oil sands bitumen, bottom of the barrel and residuum left over from refinery processes (e.g., visbreaker bottoms), and any other lower quality materials that contain a substantial quantity of high boiling hydrocarbon fractions and/or that include a significant quantity of asphaltenes that can deactivate a heterogeneous catalyst and/or cause or result in the formation of coke precursors and sediment. Examples of heavy oil feedstocks include, but are not limited to, Lloydminster heavy oil, Cold Lake bitumen, Athabasca bitumen, atmospheric tower bottoms, vacuum tower bottoms, residuum (or “resid”), resid pitch, vacuum residue (e.g., Ural VR, Arab Medium VR, Athabasca VR, Cold Lake VR, Maya VR, and Chichimene VR), deasphalted liquids obtained by solvent deasphalting, asphaltene liquids obtained as a byproduct of deasphalting, and nonvolatile liquid fractions that remain after subjecting crude oil, bitumen from tar sands, liquefied coal, oil shale, or coal tar feedstocks to distillation, hot separation, solvent extraction, and the like. By way of further example, atmospheric tower bottoms (ATB) can have a nominal boiling point of at least 343° C. (650° F.) although it is understood that the cut point can vary among refineries and be as high as 380° C. (716° F.). Vacuum tower bottoms can have a nominal boiling point of at least 524° C. (975° F.), although it is understood that the cut point can vary among refineries and be as high as 538° C. (1000° F.) or even 565° C. (1050° F.).
The term “asphaltene” shall refer to materials in a heavy oil feedstock that are typically insoluble in paraffinic solvents such as propane, butane, pentane, hexane, and heptane. Asphaltenes can include sheets of condensed ring compounds held together by hetero atoms such as sulfur, nitrogen, oxygen and metals. Asphaltenes broadly include a wide range of complex compounds having anywhere from 80 to 1200 carbon atoms, with predominating molecular weights, as determined by solution techniques, in the 1200 to 16,900 range. About 80-90% of the metals in the crude oil are contained in the asphaltene fraction which, together with a higher concentration of non-metallic hetero atoms, renders the asphaltene molecules more hydrophilic and less hydrophobic than other hydrocarbons in crude. A hypothetical asphaltene molecule structure developed by A.G. Bridge and co-workers at Chevron is depicted in
The “quality” of heavy oil is measured by at least one characteristic selected from, but not limited to: (i) boiling point; (ii) concentration of sulfur; (iii) concentration of nitrogen; (iv) concentration of metals; (v) molecular weight; (vi) hydrogen to carbon ratio; (vii) asphaltene content; and (viii) sediment forming tendency.
A “lower quality heavy oil” and/or “lower quality feedstock blend” will have at least one lower quality characteristic compared to an initial heavy oil feedstock selected from, but not limited to: (i) higher boiling point; (ii) higher concentration of sulfur; (iii) higher concentration of nitrogen; (iv) higher concentration of metals; (v) higher molecular weight (often indicated by higher density and viscosity); (vi) lower hydrogen to carbon ratio; (vii) higher asphaltene content; and (viii) greater sediment forming tendency.
The term “opportunity feedstock” refers to lower quality heavy oils and lower quality heavy oil feedstock blends having at least one lower quality characteristic compared to an initial heavy oil feedstock.
The terms “hydrocracking” and “hydroconversion” shall refer to a process whose primary purpose is to reduce the boiling range of a heavy oil feedstock and in which a substantial portion of the feedstock is converted into products with boiling ranges lower than that of the original feedstock. Hydrocracking or hydroconversion generally involves fragmentation of larger hydrocarbon molecules into smaller molecular fragments having a fewer number of carbon atoms and a higher hydrogen-to-carbon ratio. The mechanism by which hydrocracking occurs typically involves the formation of hydrocarbon free radicals during thermal fragmentation, followed by capping of the free radical ends or moieties with hydrogen. The hydrogen atoms or radicals that react with hydrocarbon free radicals during hydrocracking can be generated at or by active catalyst sites.
The term “hydrotreating” shall refer to operations whose primary purpose is to remove impurities such as sulfur, nitrogen, oxygen, halides, and trace metals from the feedstock and saturate olefins and/or stabilize hydrocarbon free radicals by reacting them with hydrogen rather than allowing them to react with themselves. The primary purpose is not to change the boiling range of the feedstock. Hydrotreating is most often carried out using a fixed bed reactor, although other hydroprocessing reactors can also be used for hydrotreating, an example of which is an ebullated bed hydrotreater.
Of course, “hydrocracking” or “hydroconversion” may also involve the removal of sulfur and nitrogen from a feedstock as well as olefin saturation and other reactions typically associated with “hydrotreating”. The terms “hydroprocessing” and “hydroconversion” shall broadly refer to both “hydrocracking” and “hydrotreating” processes, which define opposite ends of a spectrum, and everything in between along the spectrum.
The term “hydrocracking reactor” shall refer to any vessel in which hydrocracking (i.e., reducing the boiling range) of a feedstock in the presence of hydrogen and a hydrocracking catalyst is the primary purpose. Hydrocracking reactors are characterized as having an inlet port into which a heavy oil feedstock and hydrogen can be introduced, an outlet port from which an upgraded feedstock or material can be withdrawn, and sufficient thermal energy so as to form hydrocarbon free radicals in order to cause fragmentation of larger hydrocarbon molecules into smaller molecules. Examples of hydrocracking reactors include, but are not limited to, slurry phase reactors (i.e., a two phase, gas-liquid system), ebullated bed reactors (i.e., a three phase, gas-liquid-solid system), fixed bed reactors (i.e., a three-phase system that includes a liquid feed trickling downward over or flowing upward through a fixed bed of solid heterogeneous catalyst with hydrogen typically flowing cocurrently, but possibly countercurrently, to the heavy oil).
The term “hydrocracking temperature” shall refer to a minimum temperature required to cause significant hydrocracking of a heavy oil feedstock. In general, hydrocracking temperatures will preferably fall within a range of about 399° C. (750° F.) to about 460° C. (860° F.), more preferably in a range of about 418° C. (785° F.) to about 443° C. (830° F.), and most preferably in a range of about 421° C. (790° F.) to about 440° C. (825° F.).
The term “gas-liquid slurry phase hydrocracking reactor” shall refer to a hydroprocessing reactor that includes a continuous liquid phase and a gaseous dispersed phase which forms a “slurry” of gaseous bubbles within the liquid phase. The liquid phase typically comprises a hydrocarbon feedstock that may contain a low concentration of dispersed metal sulfide catalyst particles, and the gaseous phase typically comprises hydrogen gas, hydrogen sulfide, and vaporized low boiling point hydrocarbon products. The liquid phase can optionally include a hydrogen donor solvent. The term “gas-liquid-solid, 3-phase slurry hydrocracking reactor” is used when a solid catalyst is employed along with liquid and gas. The gas may contain hydrogen, hydrogen sulfide and vaporized low boiling hydrocarbon products. The term “slurry phase reactor” shall broadly refer to both type of reactors (e.g., those with dispersed metal sulfide catalyst particles, those with a micron-sized or larger particulate catalyst, and those that include both).
The terms “solid heterogeneous catalyst”, “heterogeneous catalyst” and “supported catalyst” shall refer to catalysts typically used in ebullated bed and fixed bed hydroprocessing systems, including catalysts designed primarily for hydrocracking, hydroconversion, hydrodemetallization, and/or hydrotreating. A heterogeneous catalyst typically comprises: (i) a catalyst support having a large surface area and interconnected channels or pores; and (ii) fine active catalyst particles, such as sulfides of cobalt, nickel, tungsten, and molybdenum dispersed within the channels or pores. The pores of the support are typically of limited size to maintain mechanical integrity of the heterogeneous catalyst and prevent breakdown and formation of excessive fines in the reactor. Heterogeneous catalysts can be produced as cylindrical pellets or spherical solids.
The terms “dispersed metal sulfide catalyst particles” and “dispersed catalyst” shall refer to catalyst particles having a particle size that is less than 1 μm e.g., less than about 500 nm in diameter, or less than about 250 nm in diameter, or less than about 100 nm in diameter, or less than about 50 nm in diameter, or less than about 25 nm in diameter, or less than about 10 nm in diameter, or less than about 5 nm in diameter. The term “dispersed metal sulfide catalyst particles” may include molecular or molecularly-dispersed catalyst compounds.
The term “molecularly-dispersed catalyst” shall refer to catalyst compounds that are essentially “dissolved” or dissociated from other catalyst compounds or molecules in a hydrocarbon feedstock or suitable diluent. It can include very small catalyst particles that contain a few catalyst molecules joined together (e.g., 15 molecules or less).
The terms “residual catalyst particles” shall refer to catalyst particles that remain with an upgraded material when transferred from one vessel to another (e.g., from a hydroprocessing reactor to a separator and/or other hydroprocessing reactor).
The term “conditioned feedstock” shall refer to a hydrocarbon feedstock into which a catalyst precursor has been combined and mixed sufficiently so that, upon decomposition of the catalyst precursor and formation of the active catalyst, the catalyst will comprise dispersed metal sulfide catalyst particles formed in situ within the feedstock.
The terms “upgrade”, “upgrading” and “upgraded”, when used to describe a feedstock that is being or has been subjected to hydroprocessing, or a resulting material or product, shall refer to one or more of a reduction in the molecular weight of the feedstock, a reduction in the boiling point range of the feedstock, a reduction in the concentration of asphaltenes, a reduction in the concentration of hydrocarbon free radicals, and/or a reduction in the quantity of impurities, such as sulfur, nitrogen, oxygen, halides, and metals.
The term “severity” generally refers to the amount of energy that is introduced into heavy oil during hydroprocessing and is often related to the operating temperature of the hydroprocessing reactor (i.e., higher temperature is related to higher severity; lower temperature is related to lower severity) in combination with the duration of said temperature exposure. Increased severity generally increases the quantity of conversion products produced by the hydroprocessing reactor, including both desirable products and undesirable conversion products. Desirable conversion products include hydrocarbons of reduced molecular weight, boiling point, and specific gravity, which can include end products such as naphtha, diesel, jet fuel, kerosene, wax, fuel oil, and the like. Other desirable conversion products include higher boiling hydrocarbons that can be further processed using conventional refining and/or distillation processes. Undesirable conversion products include coke, sediment, metals, and other solid materials that can deposit on hydroprocessing equipment and cause fouling, such as interior components of reactors, separators, filters, pipes, towers, and the heterogeneous catalyst. Undesirable conversion products can also refer to unconverted resid that remains after distillation, such as atmospheric tower bottoms (“ATB”) or vacuum tower bottoms (“VTB”). Minimizing undesirable conversion products reduces equipment fouling and shutdowns required to clean the equipment. Nevertheless, there may be a desirable amount of unconverted resid in order for downstream separation equipment to function properly and/or in order to provide a liquid transport medium for containing coke, sediment, metals, and other solid materials that might otherwise deposit on and foul equipment but that can be transported away by the remaining resid.
In addition to temperature, “severity” can be related to one or both of “conversion” and “throughput”. Whether increased severity involves increased conversion and/or increased or decreased throughput may depend on the quality of the heavy oil feedstock and/or the mass balance of the overall hydroprocessing system. For example, where it is desired to convert a greater quantity of feed material and/or provide a greater quantity of material to downstream equipment, increased severity may primarily involve increased throughput without necessarily increasing fractional conversion. This can include the case where resid fractions (ATB and/or VTB) are sold as fuel oil and increased conversion without increased throughput might decrease the quantity of this product. In the case where it is desired to increase the ratio of upgraded materials to resid fractions, it may be desirable to primarily increase conversion without necessarily increasing throughput. Where the quality of heavy oil introduced into the hydroprocessing reactor fluctuates, it may be desirable to selectively increase or decrease one or both of conversion and throughput to maintain a desired ratio of upgraded materials to resid fractions and/or a desired absolute quantity or quantities of end product(s) being produced.
The terms “conversion” and “fractional conversion” refer to the proportion, often expressed as a percentage, of heavy oil that is beneficially converted into lower boiling and/or lower molecular weight materials. The conversion is expressed as a percentage of the initial resid content (i.e. components with boiling point greater than a defined residue cut point) which is converted to products with boiling point less than the defined cut point. The definition of residue cut point can vary, and can nominally include 524° C. (975° F.), 538° C. (1000° F.), 565° C. (1050° F.), and the like. It can be measured by distillation analysis of feed and product streams to determine the concentration of components with boiling point greater than the defined cut point. Fractional conversion is expressed as (F−P)/F, where F is the quantity of resid in the combined feed streams, and P is the quantity in the combined product streams, where both feed and product resid content are based on the same cut point definition. The quantity of resid is most often defined based on the mass of components with boiling point greater than the defined cut point, but volumetric or molar definitions could also be used.
The term “throughput” refers to the quantity of feed material that is introduced into the hydroprocessing reactor as a function of time. It is also related to the total quantity of conversion products removed from the hydroprocessing reactor, including the combined amounts of desirable and undesirable products. Throughput can be expressed in volumetric terms, such as barrels per day, or in mass terms, such as metric tons per hour. In common usage, throughput is defined as the mass or volumetric feed rate of only the heavy oil feedstock itself (for example, vacuum tower bottoms or the like). The definition does not normally include quantities of diluents or other components that may sometimes be included in the overall feeds to a hydroconversion unit, although a definition which includes those other components could also be used.
The term “sediment” refers to solids contained in a liquid stream that can settle out. Sediments can include inorganics, coke, or insoluble asphaltenes that precipitate on cooling after conversion. Sediment in petroleum products is commonly measured using the IP-375 hot filtration test procedure for total sediment in residual fuel oils published as part of ISO 10307 and ASTM D4870. Other tests include the IP-390 sediment test and the Shell hot filtration test. Sediment is related to components of the oil that have a propensity for forming solids during processing and handling. These solid-forming components have multiple undesirable effects in a hydroconversion process, including degradation of product quality and operability problems related to fouling. It should be noted that although the strict definition of sediment is based on the measurement of solids in a sediment test, it is common for the term to be used more loosely to refer to the solids-forming components of the oil itself.
The term “fouling” refers to the formation of an undesirable phase (foulant) that interferes with processing. The foulant is normally a carbonaceous material or solid that deposits and collects within the processing equipment. Fouling can result in loss of production due to equipment shutdown, decreased performance of equipment, increased energy consumption due to the insulating effect of foulant deposits in heat exchangers or heaters, increased maintenance costs for equipment cleaning, reduced efficiency of fractionators, and reduced reactivity of heterogeneous catalyst.
Reactor 10 further includes an expanded catalyst zone 22 comprising a heterogeneous catalyst 24 that is maintained in an expanded or fluidized state against the force of gravity by upward movement of liquid hydrocarbons and gas (schematically depicted as bubbles 25) through ebullated bed reactor 10. The lower end of expanded catalyst zone 22 is defined by a distributor grid plate 26, which separates expanded catalyst zone 22 from a lower heterogeneous catalyst free zone 28 located between the bottom of ebullated bed reactor 10 and distributor grid plate 26. Distributor grid plate 26 is configured to distribute the hydrogen gas and hydrocarbons evenly across the reactor and prevents heterogeneous catalyst 24 from falling by the force of gravity into lower heterogeneous catalyst free zone 28. The upper end of the expanded catalyst zone 22 is the height at which the downward force of gravity begins to equal or exceed the uplifting force of the upwardly moving feedstock and gas through ebullated bed reactor 10 as heterogeneous catalyst 24 reaches a given level of expansion or separation. Above expanded catalyst zone 22 is an upper heterogeneous catalyst free zone 30.
Hydrocarbons and other materials within the ebullated bed reactor 10 are continuously recirculated from upper heterogeneous catalyst free zone 30 to lower heterogeneous catalyst free zone 28 by means of a recycling channel 32 positioned in the center of ebullated bed reactor 10 connected to an ebullating pump 34 at the bottom of ebullated bed reactor 10. At the top of recycling channel 32 is a funnel-shaped recycle cup 36 through which feedstock is drawn from upper heterogeneous catalyst free zone 30. Material drawn downward through recycling channel 32 enters lower catalyst free zone 28 and then passes upwardly through distributor grid plate 26 and into expanded catalyst zone 22, where it is blended with freshly added feedstock 14 and hydrogen gas 16 entering ebullated bed reactor 10 through inlet port 12. Continuously circulating blended materials upward through the ebullated bed reactor 10 advantageously maintains heterogeneous catalyst 24 in an expanded or fluidized state within expanded catalyst zone 22, minimizes channeling, controls reaction rates, and keeps heat released by the exothermic hydrogenation reactions to a safe level.
Fresh heterogeneous catalyst 24 is introduced into ebullated bed reactor 10, such as expanded catalyst zone 22, through a catalyst inlet tube 38, which passes through the top of ebullated bed reactor 10 and directly into expanded catalyst zone 22. Spent heterogeneous catalyst 24 is withdrawn from expanded catalyst zone 22 through a catalyst withdrawal tube 40 that passes from a lower end of expanded catalyst zone 22 through distributor grid plate 26 and the bottom of ebullated bed reactor 10. It will be appreciated that the catalyst withdrawal tube 40 is unable to differentiate between fully spent catalyst, partially spent but active catalyst, and freshly added catalyst such that a random distribution of heterogeneous catalyst 24 is typically withdrawn from ebullated bed reactor 10 as “spent” catalyst.
Upgraded material 20 withdrawn from ebullated bed reactor 10 can be introduced into a separator 42 (e.g., hot separator, inter-stage pressure differential separator, or distillation tower). The separator 42 separates one or more volatile fractions 46 from a non-volatile fraction 48.
Materials are continuously recirculated within reactor 110 by a recycling channel 132 connected to an ebullating pump 134 positioned outside of reactor 110. Materials are drawn through a funnel-shaped recycle cup 136 from upper catalyst free zone 130. Recycle cup 136 is spiral-shaped, which helps separate hydrogen bubbles 125 from recycles material 132 to prevent cavitation of ebullating pump 134. Recycled material 132 enters lower catalyst free zone 128, where it is blended with fresh feedstock 116 and hydrogen gas 118, and the mixture passes up through distributor grid plate 126 and into expanded catalyst zone 122. Fresh catalyst 124 is introduced into expanded catalyst zone 122 through a catalyst inlet tube 136, and spent catalyst 124 is withdrawn from expanded catalyst zone 122 through a catalyst discharge tube 140.
The main difference between the H-Oil ebullated bed reactor 110 and the LC-Fining ebullated bed reactor 10 is the location of the ebullating pump. Ebullating pump 134 in H-Oil reactor 110 is located external to the reaction chamber. The recirculating feedstock is introduced through a recirculation port 141 at the bottom of reactor 110. The recirculation port 141 includes a distributor 143, which aids in evenly distributing materials through lower catalyst free zone 128. Upgraded material 120 is shown being sent to a separator 142, which separates one or more volatile fractions 146 from a non-volatile fraction 148.
It should be understood that one or more interstage separators can optionally be interposed between first and second reactors 210a, 210b and/or second and third reactors 210b, 210c, in order to remove lower boiling fractions and gases from a non-volatile fraction containing liquid hydrocarbons and residual dispersed metal sulfide catalyst particles. It can be desirable to remove lower alkanes, such as hexanes and heptanes, which are valuable fuel products but poor solvents for asphaltenes. Removing volatile materials between multiple reactors enhances production of valuable products and increases the solubility of asphaltenes in the hydrocarbon liquid fraction fed to the downstream reactor(s). Both increase efficiency of the overall hydroprocessing system.
Upgraded material 220c from third ebullated bed reactor 210c is sent to a high temperature separator 242a, which separates volatile and non-volatile fractions. Volatile fraction 246a passes through a heat exchanger 250, which preheats hydrogen gas 216 prior to being introduced into first ebullated bed reactor 210a. The somewhat cooled volatile fraction 246a is sent to a medium temperature separator 242b, which separates a remaining volatile fraction 246b from a resulting liquid fraction 248b that forms as a result of cooling by heat exchanger 250. Remaining volatile fraction 246b is sent downstream to a low temperature separator 246c for further separation into a gaseous fraction 252c and a degassed liquid fraction 248c.
A liquid fraction 248a from high temperature separator 242a is sent together with resulting liquid fraction 248b from medium temperature separator 242b to a low pressure separator 242d, which separates a hydrogen rich gas 252d from a degassed liquid fraction 248d, which is then mixed with the degassed liquid fraction 248c from low temperature separator 242c and fractionated into products. Gaseous fraction 252c from low temperature separator 242c is purified into off gas, purge gas, and hydrogen gas 216. Hydrogen gas 216 is compressed, mixed with make-up hydrogen gas 216a, and either passed through heat exchanger 250 and introduced into first ebullated bed reactor 210a together with feedstock 216 or introduced directly into second and third ebullated bed reactors 210b and 210b.
This allows lighter, more-saturated components formed in the first two reactor stages to bypass third-stage reactor 210c. The benefits of this are (1) a reduced vapor load on the third-stage reactor, which increases the volume utilization of the third-stage reactor for converting the remaining heavy components, and (2) a reduced concentration of “anti-solvent” components (saturates) which can destabilize asphaltenes in third-stage reactor 210c.
In preferred embodiments, the hydroprocessing systems are configured and operated to promote hydrocracking reactions rather than mere hydrotreating, which is a less severe form of hydroprocessing. Hydrocracking involves the breaking of carbon-carbon molecular bonds, such as reducing the molecular weight of larger hydrocarbon molecules and/or ring opening of aromatic compounds. Hydrotreating, on the other hand, mainly involves hydrogenation of unsaturated hydrocarbons, with minimal or no breaking of carbon-carbon molecular bonds. To promote hydrocracking rather than mere hydrotreating reactions, the hydroprocessing reactor(s) are preferably operated at a temperature in a range of about 750° F. (399° C.) to about 860° F. (460° C.), more preferably in a range of about 780° F. (416° C.) to about 830° F. (443° C.), are preferably operated at a pressure in a range of about 1000 psig (6.9 MPa) to about 3000 psig (20.7 MPa), more preferably in a range of about 1500 psig (10.3 MPa) to about 2500 psig (17.2 MPa), and are preferably operated at a space velocity (e.g., Liquid Hourly Space Velocity, or LHSV, defined as the ratio of feed volume to reactor volume per hour) in a range of about 0.05 hr−1 to about 0.45 hr1, more preferably in a range of about 0.15 hr−1 to about 0.35 hr−1. The difference between hydrocracking and hydrotreating can also be expressed in terms of resid conversion (wherein hydrocracking results in the substantial conversion of higher boiling to lower boiling hydrocarbons, while hydrotreating does not). The hydroprocessing systems disclosed herein can result in a resid conversion in a range of about 40% to about 90%, preferably in a range of about 55% to about 80%. The preferred conversion range typically depends on the type of feedstock because of differences in processing difficulty between different feedstocks. Typically, conversion will be at least about 5%, preferably at least about 10% higher, compared to operating an ebullated bed reactor prior to upgrading to utilize a dual catalyst system as disclosed herein.
According to some embodiments, the heterogeneous catalyst utilized when initially operating the ebullated bed reactor at an initial condition is a commercially available catalyst that is typically used in ebullated bed reactors. To maximize efficiency, the initial reactor conditions may advantageously be with a reactor severity at which sediment formation and fouling are maintained within acceptable levels. Increasing reactor severity without upgrading the ebullated reactor to use a dual catalyst system may therefore result in excessive sediment formation and undesirable equipment fouling, which would otherwise require more frequent shutdown and cleaning of the hydroprocessing reactor and related equipment, such as pipes, towers, heaters, heterogeneous catalyst and/or separation equipment.
In order to increase reactor severity and increase the production of converted products without increasing equipment fouling and the need for more frequent shutdown and maintenance, the ebullated bed reactor is upgraded to use a dual catalyst system comprising a heterogeneous catalyst and dispersed metal sulfide catalyst particles. Operating the upgraded ebullated bed reactor with increased severity may include operating with increased conversion and/or increased throughput than when operating at the initial conditions. Both typically involve operating the upgraded reactor at an increased temperature.
In some embodiments, operating the upgraded reactor with increased reactor severity includes increasing the operating temperature of the upgraded ebullated bed reactor by nominally at least about 2.5° C., or at least about 5° C., at least about 7.5° C., or at least about 10° C., or at least about 15° C., than when operating at the initial conditions.
In some embodiments, operating the upgraded reactor with increased conversion includes increasing the conversion of the upgraded ebullated bed reactor by at least about 2.5%, or at least about 5%, at least about 7.5%, or at least about 10%, or at least about 15%, than when operating at the initial conditions.
In some embodiments, operating the upgraded reactor with increased throughput includes increasing the throughput of the upgraded ebullated bed reactor by at least about 2.5%, or at least about 5%, or at least about 10%, or at least about 15%, or at least about 20% (e.g., 24%), than when operating at the initial conditions.
In some embodiments, operating the upgraded reactor with increased conversion and throughput includes increasing the conversion of the upgraded ebullated bed reactor by at least about 2.5%, or at least about 5%, at least about 7.5%, or at least about 10%, or at least about 15%, and also increasing the throughput by at least about 2.5%, or at least about 5%, at least about 10%, or at least about 15%, or at least about 20%, than when operating at the initial conditions.
The dispersed metal sulfide catalyst particles can be generated separately and then added to the ebullated bed reactor when forming the dual catalyst system. Alternatively or in addition, at least a portion of the dispersed metal sulfide catalyst particles can be generated in situ within the ebullated bed reactor.
In some embodiments, the dispersed metal sulfide catalyst particles are advantageously formed in situ within an entirety of a heavy oil feedstock. This can be accomplished by initially mixing a catalyst precursor with an entirety of the heavy oil feedstock to form a conditioned feedstock and therefore heating the conditioned feedstock to decompose the catalyst precursor and cause or allow catalyst metal to react with sulfur in and/or added to the heavy oil to form the dispersed metal sulfide catalyst particles.
The catalyst precursor can be oil soluble and have a decomposition temperature in a range from about 100° C. (212° F.) to about 350° C. (662° F.), or in a range of about 150° C. (302° F.) to about 300° C. (572° F.), or in a range of about 175° C. (347° F.) to about 250° C. (482° F.). Example catalyst precursors include organometallic complexes or compounds, more specifically oil soluble compounds or complexes of transition metals and organic acids, having a decomposition temperature or range high enough to avoid substantial decomposition when mixed with a heavy oil feedstock under suitable mixing conditions. When mixing the catalyst precursor with a hydrocarbon oil diluent, it is advantageous to maintain the diluent at a temperature below which significant decomposition of the catalyst precursor occurs. One of skill in the art can, following the present disclosure, select a mixing temperature profile that results in intimate mixing of a selected precursor composition without substantial decomposition prior to formation of the dispersed metal sulfide catalyst particles.
Example catalyst precursors include, but are not limited to, molybdenum 2-ethylhexanoate, molybdenum octoate, molybdenum naphthanate, vanadium naphthanate, vanadium octoate, molybdenum hexacarbonyl, vanadium hexacarbonyl, and iron pentacarbonyl. Other catalyst precursors include molybdenum salts comprising a plurality of cationic molybdenum atoms and a plurality of carboxylate anions of at least 8 carbon atoms and that are at least one of (a) aromatic, (b) alicyclic, or (c) branched, unsaturated and aliphatic. By way of example, each carboxylate anion may have between 8 and 17 carbon atoms or between 11 and 15 carbon atoms. Examples of carboxylate anions that fit at least one of the foregoing categories include carboxylate anions derived from carboxylic acids selected from the group consisting of 3-cyclopentylpropionic acid, cyclohexanebutyric acid, biphenyl-2-carboxylic acid, 4-heptylbenzoic acid, 5-phenylvaleric acid, geranic acid (3,7-dimethyl-2,6-octadienoic acid), and combinations thereof.
In other embodiments, carboxylate anions for use in making oil soluble, thermally stable, molybdenum catalyst precursor compounds are derived from carboxylic acids selected from the group consisting of 3-cyclopentylpropionic acid, cyclohexanebutyric acid, biphenyl-2-carboxylic acid, 4-heptylbenzoic acid, 5-phenylvaleric acid, geranic acid (3,7-dimethyl-2,6-octadienoic acid), 10-undecenoic acid, dodecanoic acid, and combinations thereof. It has been discovered that molybdenum catalyst precursors made using carboxylate anions derived from the foregoing carboxylic acids possess improved thermal stability.
Catalyst precursors with higher thermal stability can have a first decomposition temperature higher than 210° C., higher than about 225° C., higher than about 230° C., higher than about 240° C., higher than about 275° C., or higher than about 290° C. Such catalyst precursors can have a peak decomposition temperature higher than 250° C., or higher than about 260° C., or higher than about 270° C., or higher than about 280° C., or higher than about 290° C., or higher than about 330° C.
One of skill in the art can, following the present disclosure, select a mixing temperature profile that results in intimate mixing of a selected precursor composition without substantial decomposition prior to formation of the dispersed metal sulfide catalyst particles.
Whereas it is within the scope of the invention to directly blend the catalyst precursor composition with the heavy oil feedstock, care must be taken in such cases to mix the components for a time sufficient to thoroughly blend the precursor composition within the feedstock before substantial decomposition of the precursor composition has occurred. For example, U.S. Pat. No. 5,578,197 to Cyr et al., the disclosure of which is incorporated by reference, describes a method whereby molybdenum 2-ethyl hexanoate was mixed with bitumen vacuum tower residuum for 24 hours before the resulting mixture was heated in a reaction vessel to form the catalyst compound and to effect hydrocracking (see col. 10, lines 4-43). Whereas 24-hour mixing in a testing environment may be entirely acceptable, such long mixing times may make certain industrial operations prohibitively expensive. To ensure thorough mixing of the catalyst precursor within the heavy oil prior to heating to form the active catalyst, a series of mixing steps are performed by different mixing apparatus prior to heating the conditioned feedstock. These may include one or more low shear in-line mixers, followed by one or more high shear mixers, followed by a surge vessel and pump-around system, followed by one or more multi-stage high pressure pumps used to pressurize the feed stream prior to introducing it into a hydroprocessing reactor.
In some embodiments, the conditioned feedstock is pre-heated using a heating apparatus prior to entering the hydroprocessing reactor in order to form at least a portion of the dispersed metal sulfide catalyst particles in situ within the heavy oil. In other embodiments, the conditioned feedstock is heated or further heated in the hydroprocessing reactor in order to form at least a portion of the dispersed metal sulfide catalyst particles in situ within the heavy oil.
In some embodiments, the dispersed metal sulfide catalyst particles can be formed in a multi-step process. For example, an oil soluble catalyst precursor composition can be pre-mixed with a hydrocarbon diluent to form a diluted precursor mixture. Examples of suitable hydrocarbon diluents include, but are not limited to, vacuum gas oil (which typically has a nominal boiling range of 360-524° C.) (680-975° F.), decant oil or cycle oil (which typically has a nominal boiling range of 360°-550° C.) (680-1022° F.), and gas oil (which typically has a nominal boiling range of 200°-360° C.) (392-680° F.), a portion of the heavy oil feedstock, and other hydrocarbons that nominally boil at a temperature higher than about 200° C.
The ratio of catalyst precursor to hydrocarbon oil diluent used to make the diluted precursor mixture can be in a range of about 1:500 to about 1:1, or in a range of about 1:150 to about 1:2, or in a range of about 1:100 to about 1:5 (e.g., 1:100, 1:50, 1:30, or 1:10).
The amount of catalyst metal (e.g., molybdenum) in the diluted precursor mixture is preferably in a range of about 100 ppm to about 7000 ppm by weight of the diluted precursor mixture, more preferably in a range of about 300 ppm to about 4000 ppm by weight of the diluted precursor mixture.
The catalyst precursor is advantageously mixed with the hydrocarbon diluent below a temperature at which a significant portion of the catalyst precursor composition decomposes. The mixing may be performed at temperature in a range of about 25° C. (77° F.) to about 250° C. (482° F.), or in range of about 50° C. (122° F.) to about 200° C. (392° F.), or in a range of about 75° C. (167° F.) to about 150° C. (302° F.), to form the diluted precursor mixture. The temperature at which the diluted precursor mixture is formed may depend on the decomposition temperature and/or other characteristics of the catalyst precursor that is utilized and/or characteristics of the hydrocarbon diluent, such as viscosity.
The catalyst precursor is preferably mixed with the hydrocarbon oil diluent for a time period in a range of about 0.1 second to about 5 minutes, or in a range of about 0.5 second to about 3 minutes, or in a range of about 1 second to about 1 minute. The actual mixing time is dependent, at least in part, on the temperature (i.e., which affects the viscosity of the fluids) and mixing intensity. Mixing intensity is dependent, at least in part, on the number of stages e.g., for an in-line static mixer.
Pre-blending the catalyst precursor with a hydrocarbon diluent to form a diluted precursor mixture which is then blended with the heavy oil feedstock greatly aids in thoroughly and intimately blending the catalyst precursor within the feedstock, particularly in the relatively short period of time required for large-scale industrial operations. Forming a diluted precursor mixture shortens the overall mixing time by (1) reducing or eliminating differences in solubility between a more polar catalyst precursor and a more hydrophobic heavy oil feedstock, (2) reducing or eliminating differences in rheology between the catalyst precursor and heavy oil feedstock, and/or (3) breaking up catalyst precursor molecules to form a solute within the hydrocarbon diluent that is more easily dispersed within the heavy oil feedstock.
The diluted precursor mixture is then combined with the heavy oil feedstock and mixed for a time sufficient and in a manner so as to disperse the catalyst precursor throughout the feedstock to form a conditioned feedstock in which the catalyst precursor is thoroughly mixed within the heavy oil prior to thermal decomposition and formation of the active metal sulfide catalyst particles. In order to obtain sufficient mixing of the catalyst precursor within the heavy oil feedstock, the diluted precursor mixture and heavy oil feedstock are advantageously mixed for a time period in a range of about 0.1 second to about 5 minutes, or in a range from about 0.5 second to about 3 minutes, or in a range of about 1 second to about 3 minutes. Increasing the vigorousness and/or shearing energy of the mixing process generally reduce the time required to effect thorough mixing.
Examples of mixing apparatus that can be used to effect thorough mixing of the catalyst precursor and/or diluted precursor mixture with heavy oil include, but are not limited to, high shear mixing such as mixing created in a vessel with a propeller or turbine impeller; multiple static in-line mixers; multiple static in-line mixers in combination with in-line high shear mixers; multiple static in-line mixers in combination with in-line high shear mixers followed by a surge vessel; combinations of the above followed by one or more multi-stage centrifugal pumps; and one or more multi-stage centrifugal pumps. According some embodiments, continuous rather than batch-wise mixing can be carried out using high energy pumps having multiple chambers within which the catalyst precursor composition and heavy oil feedstock are churned and mixed as part of the pumping process itself. The foregoing mixing apparatus may also be used for the pre-mixing process discussed above in which the catalyst precursor is mixed with the hydrocarbon diluent to form the catalyst precursor mixture.
In the case of heavy oil feedstocks that are solid or extremely viscous at room temperature, such feedstocks may advantageously be heated in order to soften them and create a feedstock having sufficiently low viscosity so as to allow good mixing of the oil soluble catalyst precursor into the feedstock composition. In general, decreasing the viscosity of the heavy oil feedstock will reduce the time required to effect thorough and intimate mixing of the oil soluble precursor composition within the feedstock.
The heavy oil feedstock and catalyst precursor and/or diluted precursor mixture are advantageously mixed at a temperature in a range of about 25° C. (77° F.) to about 350° C. (662° F.), or in a range of about 50° C. (122° F.) to about 300° C. (572° F.), or in a range of about 75° C. (167° F.) to about 250° C. (482° F.) to yield a conditioned feedstock.
In the case where the catalyst precursor is mixed directly with the heavy oil feedstock without first forming a diluted precursor mixture, it may be advantageous to mix the catalyst precursor and heavy oil feedstock below a temperature at which a significant portion of the catalyst precursor composition decomposes. However, in the case where the catalyst precursor is premixed with a hydrocarbon diluent to form a diluted precursor mixture, which is thereafter mixed with the heavy oil feedstock, it may be permissible for the heavy oil feedstock to be at or above the decomposition temperature of the catalyst precursor. That is because the hydrocarbon diluent shields the individual catalyst precursor molecules and prevents them from agglomerating to form larger particles, temporarily insulates the catalyst precursor molecules from heat from the heavy oil during mixing, and facilitates dispersion of the catalyst precursor molecules sufficiently quickly throughout the heavy oil feedstock before decomposing to liberate metal. In addition, additional heating of the feedstock may be necessary to liberate hydrogen sulfide from sulfur-bearing molecules in the heavy oil to form the metal sulfide catalyst particles. In this way, progressive dilution of the catalyst precursor permits a high level of dispersion within the heavy oil feedstock, resulting in the formation of highly dispersed metal sulfide catalyst particles, even where the feedstock is at a temperature above the decomposition temperature of the catalyst precursor.
After the catalyst precursor has been well-mixed throughout the heavy oil to yield a conditioned feedstock, this composition is then heated to cause decomposition of the catalyst precursor to liberate catalyst metal therefrom, cause or allow it to react with sulfur within and/or added to the heavy oil, and form the active metal sulfide catalyst particles. Metal from the catalyst precursor may initially form a metal oxide, which then reacts with sulfur in the heavy oil to yield a metal sulfide compound that forms the final active catalyst. In the case where the heavy oil feedstock includes sufficient or excess sulfur, the final activated catalyst may be formed in situ by heating the heavy oil feedstock to a temperature sufficient to liberate sulfur therefrom. In some cases, sulfur may be liberated at the same temperature that the precursor composition decomposes. In other cases, further heating to a higher temperature may be required.
If the catalyst precursor is thoroughly mixed throughout the heavy oil, at least a substantial portion of the liberated metal ions will be sufficiently sheltered or shielded from other metal ions so that they can form a molecularly-dispersed catalyst upon reacting with sulfur to form the metal sulfide compound. Under some circumstances, minor agglomeration may occur, yielding colloidal-sized catalyst particles. However, it is believed that taking care to thoroughly mix the catalyst precursor throughout the feedstock prior to thermal decomposition of the catalyst precursor may yield individual catalyst molecules rather than colloidal particles. Simply blending, while failing to sufficiently mix, the catalyst precursor with the feedstock typically causes formation of large agglomerated metal sulfide compounds that are micron-sized or larger.
In order to form dispersed metal sulfide catalyst particles, the conditioned feedstock is heated to a temperature in a range of about 275° C. (527° F.) to about 450° C. (842° F.), or in a range of about 310° C. (590° F.) to about 430° C. (806° F.), or in a range of about 330° C. (626° F.) to about 410° C. (770° F.).
The initial concentration of catalyst metal provided by dispersed metal sulfide catalyst particles can be in a range of about 1 ppm to about 500 ppm by weight of the heavy oil feedstock, or in a range of about 5 ppm to about 300 ppm, or in a range of about 10 ppm to about 100 ppm. The catalyst may become more concentrated as volatile fractions are removed from a resid fraction.
In the case where the heavy oil feedstock includes a significant quantity of asphaltene molecules, the dispersed metal sulfide catalyst particles may preferentially associate with, or remain in close proximity to, the asphaltene molecules. Asphaltene molecules can have a greater affinity for the metal sulfide catalyst particles since asphaltene molecules are generally more hydrophilic and less hydrophobic than other hydrocarbons contained within heavy oil. Because the metal sulfide catalyst particles tend to be very hydrophilic, the individual particles or molecules will tend to migrate toward more hydrophilic moieties or molecules within the heavy oil feedstock.
While the highly polar nature of metal sulfide catalyst particles causes or allows them to associate with asphaltene molecules, it is the general incompatibility between the highly polar catalyst compounds and hydrophobic heavy oil that necessitates the aforementioned intimate or thorough mixing of catalyst precursor composition within the heavy oil prior to decomposition and formation of the active catalyst particles. Because metal catalyst compounds are highly polar, they cannot be effectively dispersed within heavy oil if added directly thereto. In practical terms, forming smaller active catalyst particles results in a greater number of catalyst particles that provide more evenly distributed catalyst sites throughout the heavy oil.
The conditioned feedstock from surge vessel 412 is pressurized by one or more pumps 416, passed through a pre-heater 418, and fed into ebullated bed reactor 430 together with pressurized hydrogen gas 420 through an inlet port 436 located at or near the bottom of ebullated bed reactor 430. Heavy oil material 426 in ebullated bed reactor 430 contains dispersed metal sulfide catalyst particles, schematically depicted as catalyst particles 424.
Heavy oil feedstock 408 may comprise any desired fossil fuel feedstock and/or fraction thereof including, but not limited to, one or more of heavy crude, oil sands bitumen, bottom of the barrel fractions from crude oil, atmospheric tower bottoms, vacuum tower bottoms, coal tar, liquefied coal, and other resid fractions. In some embodiments, heavy oil feedstock 408 can include a significant fraction of high boiling point hydrocarbons (i.e., nominally at or above 343° C. (650° F.), more particularly nominally at or above about 524° C. (975° F.)) and/or asphaltenes. Asphaltenes are complex hydrocarbon molecules that include a relatively low ratio of hydrogen to carbon that is the result of a substantial number of condensed aromatic and naphthenic rings with paraffinic side chains (See
Ebullated bed reactor 430 further includes an expanded catalyst zone 442 comprising a heterogeneous catalyst 444. A lower heterogeneous catalyst free zone 448 is located below expanded catalyst zone 442, and an upper heterogeneous catalyst free zone 450 is located above expanded catalyst zone 442. Dispersed metal sulfide catalyst particles 424 are dispersed throughout material 426 within ebullated bed reactor 430, including expanded catalyst zone 442, heterogeneous catalyst free zones 448, 450, 452 thereby being available to promote upgrading reactions within what constituted catalyst free zones in the ebullated bed reactor prior to being upgraded to include the dual catalyst system.
To promote hydrocracking rather than mere hydrotreating reactions, the hydroprocessing reactor(s) are preferably operated at a temperature in a range of about 750° F. (399° C.) to about 860° F. (460° C.), more preferably in a range of about 780° F. (416° C.) to about 830° F. (443° C.), are preferably operated at a pressure in a range of about 1000 psig (6.9 MPa) to about 3000 psig (20.7 MPa), more preferably in a range of about 1500 psig (10.3 MPa) to about 2500 psig (17.2 MPa), and are preferably operated at a space velocity (LHSV) in a range of about 0.05 hr−1 to about 0.45 hr−1, more preferably in a range of about 0.15 hr−1 to about 0.35 hr−1. The difference between hydrocracking and hydrotreating can also be expressed in terms of resid conversion (wherein hydrocracking results in the substantial conversion of higher boiling to lower boiling hydrocarbons, while hydrotreating does not). The hydroprocessing systems disclosed herein can result in a resid conversion in a range of about 40% to about 90%, preferably in a range of about 55% to about 80%. The preferred conversion range typically depends on the type of feedstock because of differences in processing difficulty between different feedstocks. Typically, conversion will be at least about 5%, preferably at least about 10% higher, compared to operating an ebullated bed reactor prior to upgrading to utilize a dual catalyst system as disclosed herein.
Material 426 in ebullated bed reactor 430 is continuously recirculated from upper heterogeneous catalyst free zone 450 to lower heterogeneous catalyst free zone 448 by means of a recycling channel 452 connected to an ebullating pump 454. At the top of recycling channel 452 is a funnel-shaped recycle cup 456 through which material 426 is drawn from upper heterogeneous catalyst free zone 450. Recycled material 426 is blended with fresh conditioned feedstock 411 and hydrogen gas 420.
Fresh heterogeneous catalyst 444 is introduced into ebullated bed reactor 430 through a catalyst inlet tube 458, and spent heterogeneous catalyst 444 is withdrawn through a catalyst withdrawal tube 460. Whereas the catalyst withdrawal tube 460 is unable to differentiate between fully spent catalyst, partially spent but active catalyst, and fresh catalyst, the existence of dispersed metal sulfide catalyst particles 424 provides additional catalytic activity, within expanded catalyst zone 442, recycle channel 452, and lower and upper heterogeneous catalyst free zones 448, 450. The addition of hydrogen to hydrocarbons outside of heterogeneous catalyst 444 minimizes formation of sediment and coke precursors, which are often responsible for deactivating the heterogeneous catalyst.
Ebullated bed reactor 430 further includes an outlet port 438 at or near the top through which converted material 440 is withdrawn. Converted material 440 is introduced into hot separator or distillation tower 404. Hot separator or distillation tower 404 separates one or more volatile fractions 405, which is/are withdrawn from the top of hot separator 404, from a resid fraction 407, which is withdrawn from a bottom of hot separator or distillation tower 404. Resid fraction 407 contains residual metal sulfide catalyst particles, schematically depicted as catalyst particles 424. If desired, at least a portion of resid fraction 407 can be recycled back to ebullated bed reactor 430 in order to form part of the feed material and to supply additional metal sulfide catalyst particles. Alternatively, resid fraction 407 can be further processed using downstream processing equipment, such as another ebullated bed reactor. In that case, separator 404 can be an interstage separator.
In some embodiments, operating the upgraded ebullated bed reactor at a higher reactor severity and an increased rate of production of converted products while using the dual catalyst system results in a rate of equipment fouling that is equal to or less than when initially operating the ebullated bed reactor.
For example, the rate of equipment fouling when operating the upgraded ebullated bed reactor using the dual catalyst system may result in a frequency of heat exchanger shutdowns for cleanout that is equal to or less than when initially operating the ebullated bed reactor.
In addition or alternatively, the rate of equipment fouling when operating the upgraded ebullated bed reactor using the dual catalyst system may result in a frequency of atmospheric and/or vacuum distillation tower shutdowns for cleanout that is equal or less than when initially operating the ebullated bed reactor.
In addition or alternatively, the rate of fouling when operating of the upgraded ebullated bed reactor using the dual catalyst system may result in a frequency of changes or cleaning of filters and strainers that is equal or less than when initially operating the ebullated bed reactor.
In addition or alternatively, the rate of fouling when operating of the upgraded ebullated bed reactor using the dual catalyst system may result in a frequency of switches to spare heat exchangers that is equal or less than when initially operating the ebullated bed reactor.
In addition or alternatively, the rate of fouling when operating of the upgraded ebullated bed reactor using the dual catalyst system may result in a reduced rate of decreasing skin temperatures in equipment selected from one or more of heat exchangers, separators, or distillation towers than when initially operating the ebullated bed reactor.
In addition or alternatively, the rate of fouling when operating of the upgraded ebullated bed reactor using the dual catalyst system may result in a reduced rate of increasing furnace tube metal temperatures than when initially operating the ebullated bed reactor.
In addition or alternatively, the rate of fouling when operating of the upgraded ebullated bed reactor using the dual catalyst system may result in a reduced rate of increasing calculated fouling resistance factors for heat exchangers than when initially operating the ebullated bed reactor.
In some embodiments, operating the upgraded ebullated bed reactor while using the dual catalyst system may result in a rate of sediment production that is equal to or less than when initially operating the ebullated bed reactor. In some embodiments, the rate of sediment production can be based on a measurement of sediment in one or more of: (1) an atmospheric tower bottoms product; (2) a vacuum tower bottoms product; (3) product from a hot low pressure separator; or (4) fuel oil product before or after addition of cutter stocks.
In some embodiments, operating the upgraded ebullated bed reactor while using the dual catalyst system may result in a product sediment concentration that is equal or less than when initially operating the ebullated bed reactor. In some embodiments, the product sediment concentration can be based on a measurement of sediment in one or more of (1) an atmospheric residue product cut and/or an atmospheric tower bottoms product; (2) a vacuum residue product cut and/or a vacuum tower bottoms product; (3) material fed to an atmospheric tower; (4) product from a hot low pressure separator; or (5) fuel oil product before or after addition of one or more cutter stocks.
The following test studies demonstrate the effects and advantages of upgrading an ebullated bed reactor to use a dual catalyst system comprised of a heterogeneous catalyst and dispersed metal sulfide catalyst particles when hydroprocessing heavy oil. The pilot plant used for this test was designed according to
For the following test studies, a heavy vacuum gas oil was used as the hydrocarbon diluent. The precursor mixture was prepared by mixing an amount of catalyst precursor with an amount of hydrocarbon diluent to form a catalyst precursor mixture and then mixing an amount of the catalyst precursor mixture with an amount of heavy oil feedstock to achieve the target loading of dispersed catalyst in the conditioned feedstock. As a specific illustration, for one test study with a target loading of 30 ppm dispersed metal sulfide catalyst in the conditioned feedstock (where the loading is expressed based on metal concentration), the catalyst precursor mixture was prepared with a 3000 ppm concentration of metal.
The feedstocks and operating conditions for the actual tests are more particularly identified below. The heterogeneous catalyst was a commercially available catalyst commonly used in ebullated reactors. Note that for comparative test studies for which no dispersed metal sulfide catalyst was used, the hydrocarbon diluent (heavy vacuum gas oil) was added to the heavy oil feedstock in the same proportion as when using a diluted precursor mixture. This ensured that the background composition was the same between tests using the dual catalyst system and those using only the heterogeneous (ebullated bed) catalyst, thereby allowing test results to be compared directly.
Pilot plant 500 more particularly included a high shear mixing vessel 502 for blending a precursor mixture comprised of a hydrocarbon diluent and catalyst precursor (e.g., molybdenum 2-ethylhexanoate) with a heavy oil feedstock (collectively depicted as 501) to form a conditioned feedstock. Proper blending can be achieved by first pre-blending the catalyst precursor with a hydrocarbon diluent to form a precursor mixture.
The conditioned feedstock is recirculated out and back into the mixing vessel 502 by a pump 504, similar to a surge vessel and pump-around. A high precision positive displacement pump 506 draws the conditioned feedstock from the recirculation loop and pressurizes it to the reactor pressure. Hydrogen gas 508 is fed into the pressurized feedstock and the resulting mixture is passed through a pre-heater 510 prior to being introduced into first ebullated bed reactor 512. The pre-heater 510 can cause at least a portion of the catalyst precursor within the conditioned feedstock to decompose and form active catalyst particles in situ within the feedstock.
Each ebullated bed reactor 512, 512′ can have a nominal interior volume of about 3000 ml and include a mesh wire guard 514 to keep the heterogeneous catalyst within the reactor. Each reactor is also equipped with a recycle line and recycle pump 513, 513′ which provides the required flow velocity in the reactor to expand the heterogeneous catalyst bed. The combined volume of both reactors and their respective recycle lines, all of which are maintained at the specified reactor temperature, can be considered to be the thermal reaction volume of the system and can be used as the basis for calculation of the Liquid Hourly Space Velocity (LHSV). For these examples, “LHSV” is defined as the volume of vacuum residue feedstock fed to the reactor per hour divided by the thermal reaction volume.
A settled height of catalyst in each reactor is schematically indicated by a lower dotted line 516, and the expanded catalyst bed during use is schematically indicated by an upper dotted line 518. A recirculating pump 513 is used to recirculate the material being processed from the top to the bottom of reactor 512 to maintain steady upward flow of material and expansion of the catalyst bed.
Upgraded material from first reactor 512 is transferred together with supplemental hydrogen 520 into second reactor 512′ for further hydroprocessing. A second recirculating pump 513′ is used to recirculate the material being processed from the top to the bottom of second reactor 512′ to maintain steady upward flow of material and expansion of the catalyst bed.
The further upgraded material from second reactor 512′ is introduced into a hot separator 522 to separate low-boiling hydrocarbon product vapors and gases 524 from a liquid fraction 526 comprised of unconverted heavy oil. The hydrocarbon product vapors and gases 524 are cooled and pass into a cold separator 528, where they are separated into gases 530 and converted hydrocarbon products, which are recovered as separator overheads 532. The liquid fraction 526 from hot separator 522 is recovered as separator bottoms 534, which can be used for analysis.
Examples 1-4 were conducted in the abovementioned pilot plant and tested the ability of an upgraded ebullated bed reactor that employed a dual catalyst system to operate at substantially higher conversion at equal feed rate (throughput) while maintaining or reducing formation of sediment. The increased conversion included higher resid conversion, C7 asphaltene conversion, and micro carbon residue (MCR) conversion. The heavy oil feedstock utilized in this study was Ural vacuum resid (VR). As described above, a conditioned feedstock was prepared by mixing an amount of catalyst precursor mixture with an amount of heavy oil feedstock to a final conditioned feedstock that contained the required amount of dispersed catalyst. The exception to this were tests for which no dispersed catalyst was used, in which case heavy vacuum gas oil was substituted for the catalyst precursor mixture at the same proportion. The conditioned feedstock was fed into the pilot plant system of
Examples 1 and 2 utilized a heterogeneous catalyst to simulate an ebullated bed reactor prior to being upgraded to employ a dual catalyst system according to the invention. Examples 3 and 4 utilized a dual catalyst system comprised of the same heterogeneous catalyst of Examples 1 and 2 and also dispersed molybdenum sulfide catalyst particles. The concentration of dispersed molybdenum sulfide catalyst particles in the feedstock was measured as concentration in parts per million (ppm) of molybdenum metal (Mo) provided by the dispersed catalyst. The feedstock of Examples 1 and 2 included no dispersed catalyst (0 ppm Mo), the feedstock of Example 3 included dispersed catalyst at a concentration of 30 ppm Mo, and the feedstock of Example 4 included dispersed catalyst at a higher concentration of 50 ppm Mo.
Example 1 was the baseline test in which Ural VR was hydroprocessed at a temperature of 789° F. (421° C.) and a resid conversion of 60.0%. In Example 2, the temperature was increased to 801° F. (427° C.) and resid conversion (based on 1000° F.+, %) was increased to 67.7%. This resulted in a substantial increase in product IP-375 sediment (separator bottoms basis, wt. %) of 0.78% to 1.22%, product IP-375 sediment (feed oil basis, wt. %) of 0.67% to 0.98%, a C7 asphaltene conversion of 40.6% to 43.0%, and MCR conversion of 49.3% to 51.9%. This indicates that the heterogeneous catalyst used by itself in Examples 1 and 2 could not withstand an increase in temperature and conversion without a substantial increase in sediment formation.
In Example 3, which utilized the dual catalyst system, including dispersed catalyst (providing 30 ppm Mo), reactor temperature was increased to 801° F. (427° C.) and resid conversion was increased to 67.0%. Feed rate was increased slightly from 0.24 to 0.25 (LHSV, vol. feed/vol. reactor/hour). Even at higher temperature, resid conversion, and feed rate, there was a slight decrease in product IP-375 sediment (separator bottoms basis, wt. %) of 0.78% to 0.76%, a more substantial decrease in product IP-375 sediment (feed oil basis, wt. %) of 0.67% to 0.61%. In addition to increased resid conversion, the C7 asphaltene conversion was increased from 40.6% to 46.9%, and MCR conversion was increased from 49.3% to 55.2%.
The dual catalyst system of Example 3 also substantially outperformed the heterogeneous catalyst used by itself in Example 2 by a wide margin, including further increasing C7 asphaltene conversion from 43.0% to 46.9% and MCR conversion from 51.9% to 55.2%, while substantially decreasing product IP-375 sediment (separator bottoms basis, wt. %) from 1.22% to 0.76%, and product IP-375 sediment (feed oil basis, wt. %) from 0.98% to 0.61%.
In Example 4, which utilized the dual catalyst system, including dispersed catalyst (providing 50 ppm Mo), reactor temperature was 801° F. (427° C.), conversion was 65.9%, and feed rate was 0.25 (LHSV, vol. feed/vol. reactor/hour). Compared to Example 1, there was a substantial decrease in product IP-375 sediment (separator bottoms basis, wt. %) of 0.78% to 0.54%, a substantial decrease in product IP-375 sediment (feed oil basis, wt. %) of 0.67% to 0.45%. In addition, the C7 asphaltene conversion was increased from 40.6% to 46.9%, and MCR conversion was increased from 49.3% to 54.8%. This indicates that the dual catalyst system of Example 4 also substantially outperformed the heterogeneous catalyst used by itself in Example 2 by an even wider margin, including further increasing C7 asphaltene conversion from 43.0 to 46.9% and MCR conversion from 51.9% to 54.8%, while decreasing product IP-375 sediment (separator bottoms basis, wt. %) from 1.22% to 0.54%, and product IP-375 sediment (feed oil basis, wt. %) from 0.98% to 0.45%.
Examples 3 and 4 clearly demonstrated the ability of a dual catalyst system in an upgraded ebullated hydroprocessing reactor to permit increased reactor severity, including increased operating temperature, resid conversion, C7 asphaltene conversion, and MCR conversion, and equal feed rate (throughput) while substantially reducing sediment production, compared to an ebullated bed reactor using only a heterogeneous catalyst.
Examples 5-8 were conducted in the aforementioned pilot plant and also tested the ability of an upgraded ebullated bed reactor that employed a dual catalyst system to operate at substantially higher conversion at equal feed rate (throughput) while maintaining or reducing formation of sediment. The increased conversion included higher resid conversion, C7 asphaltene conversion, and micro carbon residue (MCR) conversion. The heavy oil feedstock utilized in this study was Arab Medium vacuum resid (VR). Relevant process conditions and results are set forth in Table 2.
It is noted that the sediment data for Examples 5 and 6 may conceptually have the wrong directional trend for sediment production (i.e., lower sediment at higher resid conversion while using the same heterogeneous catalyst and no dispersed catalyst). Nevertheless, the results comparing Examples 6-8 demonstrated a clear improvement when using the dual catalyst system.
Examples 5 and 6 utilized a heterogeneous catalyst to simulate an ebullated bed reactor prior to being upgraded to employ a dual catalyst system according to the invention. Examples 7 and 8 utilized a dual catalyst system comprised of the same heterogeneous catalyst of Examples 5 and 6 and dispersed molybdenum sulfide catalyst particles. The concentration of dispersed molybdenum sulfide catalyst particles in the feedstock was measured as concentration in parts per million (ppm) of molybdenum metal (Mo) provided by the dispersed catalyst. The feedstock of Examples 5 and 6 included no dispersed catalyst (0 ppm Mo); the feedstock of Example 7 included dispersed catalyst (30 ppm Mo), and the feedstock of Example 8 included dispersed catalyst (50 ppm Mo).
Example 5 was the baseline test in which Arab Medium VR was hydroprocessed at a temperature of 803° F. (428° C.) and a resid conversion of 73.2%. In Example 6, the temperature was increased to 815° F. (435° C.) and resid conversion (based on 1000° F.+, %) was increased to 81.4%. The product IP-375 sediment (separator bottoms basis, wt. %) decreased from 1.40% to 0.91%, product IP-375 sediment (feed oil basis, wt. %) decreased from 1.05% to 0.61%, C7 asphaltene conversion increased from 55.8% to 65.9%, and MCR conversion increased from 47.2% to 55.2%. For purposes of comparing the effect of the dual catalyst system of Examples 7 and 8, either Example 5 and 6 can be used. However, the most direct comparison is to the results in Example 6, which was conducted at a resid conversion essentially the same as for Examples 7 and 8.
In Example 7, which utilized dispersed catalyst particles (providing 30 ppm Mo), reactor temperature was increased from to 803° F. (428° C.) in Example 5 to 815° F. (435° C.) and resid conversion was increased to from 73.2% in Example 5 to 79.9%. Feed rate was maintained at 0.25 (LHSV, vol. feed/vol. reactor/hour). Even at higher temperature, conversion and feed rate, there was a decrease in product IP-375 sediment (separator bottoms basis, wt. %) from 1.40% to 0.68%, a decrease in product IP-375 sediment (feed oil basis, wt. %) of 1.05% to 0.49%. In addition to increased resid conversion, the C7 asphaltene conversion was increased from 55.8% to 72.9%, and MCR conversion was increased from 47.2% to 57.7%.
The dual catalyst system of Example 7 also substantially outperformed the heterogeneous catalyst used by itself in Example 6 by a wide margin, including further increasing C7 asphaltene conversion from 65.9% to 72.9% and MCR conversion from 55.2% to 57.7%, while substantially decreasing product IP-375 sediment (separator bottoms basis, wt. %) from 0.91% to 0.68%, and product IP-375 sediment (feed oil basis, wt. %) from 0.61% to 0.49%.
In Example 8, which utilized dispersed catalyst particles (providing 50 ppm Mo), reactor temperature was 815° F. (435° C.), conversion was 80.8%, and feed rate was 0.25 (LHSV, vol. feed/vol. reactor/hour). Compared to Example 5, there was a substantial decrease in product IP-375 sediment (separator bottoms basis, wt. %) from 1.40% to 0.43%, a substantial decrease in product IP-375 sediment (feed oil basis, wt. %) of 1.05% to 0.31%. In addition, the C7 asphaltene conversion was increased from 55.8% to 76.0%, and MCR conversion was increased from 47.2% to 61.8%.
The dual catalyst system of Example 8 also substantially outperformed the heterogeneous catalyst used by itself in Example 6, including further increasing C7 asphaltene conversion from 65.9 to 76.0% and MCR conversion from 55.2% to 61.8%, while decreasing product IP-375 sediment (separator bottoms basis, wt. %) from 0.91% to 0.43%, and product IP-375 sediment (feed oil basis, wt. %) from 0.61% to 0.31%.
Examples 7 and 8 clearly demonstrated the ability of a dual catalyst system in an upgraded ebullated bed hydroprocessing reactor to permit increased reactor severity, including increased operating temperature, resid conversion, C7 asphaltene conversion, and MCR conversion, and equal feed rate (throughput) while substantially reducing sediment production, compared to an ebullated bed reactor using only a heterogeneous catalyst.
Examples 9-13 are commercial results showing the ability of an upgraded ebullated bed reactor that employed a dual catalyst system to permit substantially higher conversion at equal feed rate (throughput) while maintaining or reducing formation of sediment. The increased conversion included higher resid conversion, C7 asphaltene conversion, and micro carbon residue (MCR) conversion. The heavy oil feedstock utilized in this study was Ural vacuum resid (VR). The data in this study only shows relative rather than absolute results to maintain customer confidentiality. Relevant process conditions and results are set forth in Table 1.
Example 9 utilized a heterogeneous catalyst in an ebullated bed reactor prior to being upgraded to employ a dual catalyst system according to the invention. Examples 10-13 utilized a dual catalyst system comprised of the same heterogeneous catalyst of Example 9 and dispersed molybdenum sulfide catalyst particles. The concentration of dispersed molybdenum sulfide catalyst particles in the feedstock was measured as concentration in parts per million (ppm) of molybdenum metal (Mo) provided by the dispersed catalyst. The feedstock of Example 9 included no dispersed catalyst (0 ppm Mo); the feedstocks of Examples 10-13 included dispersed catalyst (32 ppm Mo).
Example 9 was the baseline test in which Ural VR was hydroprocessed at a base temperature (Tbase), base feed rate (LHSVbase), a base resid conversion (Convbase), base sediment formation (Sedbase), base C7 conversion (C7 base), and base MCR conversion (MCRbase).
In Example 10, the temperature (Tbase) and feed rate (LHSVbase) were the same as in Example 9. Including dispersed catalyst resulted in a slight decrease in resid conversion of 1.3% compared to the base resid conversion (Convbase−1.3%), a decrease in product IP-375 sediment (separator bottoms basis, wt. %) of 0.12% (Sedbase−0.12%), a decrease in product IP-375 sediment (feed oil basis, wt. %) of 0.02% (Sedbase−0.02%), an increase in C7 asphaltene conversion of 18% (C7 base+18%), and no change in MCR conversion (MCRbase). This indicates that by simply upgrading the ebullated bed reactor to include the dual catalyst system (Example 10) instead of the heterogeneous catalyst used by itself (Example 9), C7 asphaltene conversion was increased substantially while sediment formation decreased. Even though resid conversion decreased slightly, the far more important statistic is the increase in C7 asphaltene conversion since that is the component most responsible for coke formation and equipment fouling.
In Example 11, the temperature (Tbase) was increased by 4° C. (Tbase+4° C.) compared to Example 9 and the feed rate (LHSVbase) was the same. This resulted in increased resid conversion of 2.7% (Convbase+2.7%), a decrease in product IP-375 sediment (separator bottoms basis, wt. %) of 0.09% (Sedbase−0.09%), a decrease in product IP-375 sediment (feed oil basis, wt. %) of 0.05% (Sedbase−0.05%), an increase in C7 asphaltene conversion of 25% (C7 base+25%), and an increase in MCR conversion of 2% (MCRbase+2%). This indicates that upgrading the ebullated bed reactor to include the dual catalyst system instead of the heterogeneous catalyst used by itself increased resid conversion, substantially increased C7 asphaltene conversion, increased MCR conversion, while decreasing sediment formation. While resid conversion increased slightly, the far more important statistic is the substantially higher increase in C7 asphaltene conversion.
In Example 12, the temperature (Tbase) was increased by 6° C. (Tbase+6° C.) compared to Example 9 and the feed rate (LHSVbase) was the same. This resulted in a substantially higher resid conversion of 6.3% (Convbase+6.3%), a decrease in product IP-375 sediment (separator bottoms basis, wt. %) of 0.06% (Sedbase−0.06%), a decrease in product IP-375 sediment (feed oil basis, wt. %) of 0.05% (Sedbase−0.05%), an increase in C7 asphaltene conversion of 25% (C7 base+25%), and an increase in MCR conversion of 3% (MCRbase+3%). This indicates that upgrading the ebullated bed reactor to include the dual catalyst system instead of the heterogeneous catalyst used by itself substantially increased resid conversion, C7 asphaltene conversion, increased MCR conversion, while decreasing sediment formation.
In Example 13, the temperature (Tbase) was increased by 9° C. (Tbase+9° C.) compared to Example 9 and the feed rate (LHSVbase) was the same. This resulted in a substantially higher resid conversion of 10.4% (Convbase+10.4%), a decrease in product IP-375 sediment (separator bottoms basis, wt. %) of 0.07% (Sedbase−0.07%), a decrease in product IP-375 sediment (feed oil basis, wt. %) of 0.07% (Sedbase−0.07%), an increase in C7 asphaltene conversion of 18% (C7 base+18%), and an increase in MCR conversion of 4% (MCRbase+4%). This indicates that upgrading the ebullated bed reactor to include the dual catalyst system instead of the heterogeneous catalyst used by itself substantially increased resid conversion, C7 asphaltene conversion, and MCR conversion, while decreasing sediment formation.
Examples 10-13 clearly demonstrated the ability of a dual catalyst system in an upgraded ebullated hydroprocessing reactor to permit increased reactor severity, including increased operating temperature, resid conversion, C7 asphaltene conversion, and MCR conversion, and equal feed rate (throughput) while substantially reducing sediment production, compared to an ebullated bed reactor using only a heterogeneous catalyst.
In addition to the data shown in Table 3,
Examples 14-16 were conducted in the aforementioned pilot plant and tested the ability of an upgraded ebullated bed reactor that employed a dual catalyst system to operate at substantially higher feed rate (throughput) at equal resid conversion while maintaining or reducing formation of sediment. The heavy oil feedstock utilized in this study was Arab medium vacuum resid (VR). Relevant process conditions and results are set forth in Table 4.
Examples 14 and 15 utilized a heterogeneous catalyst to simulate an ebullated bed reactor prior to being upgraded to employ a dual catalyst system according to the invention. Example 16 utilized a dual catalyst system comprised of the same heterogeneous catalyst of Examples 14 and 15 and dispersed molybdenum sulfide catalyst particles. The concentration of dispersed molybdenum sulfide catalyst particles in the feedstock was measured as concentration in parts per million (ppm) of molybdenum metal (Mo) provided by the dispersed catalyst. The feedstock of Examples 14 and 15 included no dispersed catalyst (0 ppm Mo); the feedstock of Example 16 included dispersed catalyst (30 ppm Mo).
Example 14 was the baseline test in which Arab Medium VR was hydroprocessed at a temperature of 788° F. (420° C.) and a resid conversion of 62%. In Example 15, the temperature was increased to 800° F. (427° C.), resid conversion was maintained at 62%, and feed rate (LHSV, vol. feed/vol. reactor/hour) was increased to 0.33. This resulted in a substantial increase in product IP-375 sediment (separator bottoms basis, wt. %) from 0.37% to 0.57%, increased product IP-375 sediment (feed oil basis, wt. %) from 0.30% to 0.44%, a C7 substantial decrease in asphaltene conversion of 58.0% to 48.0%, and a decrease in MCR conversion from 58.5% to 53.5%. This indicates that the heterogeneous catalyst used by itself in Examples 14 and 15 could not withstand an increase in temperature and feed rate without a substantial increase in sediment formation.
In Example 16, which utilized dispersed catalyst particles (providing 30 ppm Mo), reactor temperature was increased to 803° F. (428° C.), resid conversion was maintained at 62%, and feed rate was increased from 0.24 to 0.3 (LHSV, vol. feed/vol. reactor/hour). Even at higher temperature and feed rate, while maintaining the same resid conversion, there was a substantial decrease in product IP-375 sediment (separator bottoms basis, wt. %) from 0.37% to 0.10%, a substantial decrease in product IP-375 sediment (feed oil basis, wt. %) from 0.30% to 0.08%. In addition, the C7 asphaltene conversion increased from 58.0% to 59.5% and the MCR conversion decreased from 58.5% to 57.0%.
The dual catalyst system of Example 16 also substantially outperformed the heterogeneous catalyst in Example 15 by a wide margin, including substantially decreasing product IP-375 sediment (separator bottoms basis, wt. %) from 0.57% to 0.10%, substantially decreasing product IP-375 sediment (feed oil basis, wt. %) from 0.44% to 0.08%, substantially increasing C7 asphaltene conversion from 48.0% to 59.5%, and increasing MCR conversion from 53.5% to 57.0%.
In addition to the data shown in Table 3,
The present invention may be embodied in other specific forms without departing from its spirit or essential characteristics. The described embodiments are to be considered in all respects only as illustrative and not restrictive. The scope of the invention is, therefore, indicated by the appended claims rather than by the foregoing description. All changes which come within the meaning and range of equivalency of the claims are to be embraced within their scope.
This application claims the benefit of U.S. Provisional Patent Application No. 62/222,073, filed Sep. 22, 2016, the disclosure of which is incorporated herein in its entirety.
| Number | Date | Country | |
|---|---|---|---|
| 62222073 | Sep 2015 | US |
