The present disclosure generally relates to upgrading difficult to process heavy-oil. In particular, the disclosure relates to upgrading heavy oil and other high carbon content materials by using one or more processes that directly incorporate lighter hydrocarbons into high molecular weight, low hydrogen content hydrocarbons such as thermally processed heavy oil products.
Heavy oil can be upgraded and ultimately refined into various commercially valuable products, including fuels and chemicals. The goals of conventional heavy-oil upgrading systems and processes, each of which systems are also referred to as a heavy-oil thermal processor, include: removing impurities such as nitrogen and sulfur; hydrogenating (saturating) olefins; opening aromatic structures; and, cracking long chain, high molecular-weight compounds into shorter chain, lower molecular-weight compounds.
An initial step in upgrading heavy oil is typically a low-temperature distillation process, such as atmospheric-pressure distillation, which separates valuable precursor materials from heavier materials referred to as Atmospheric Tower Distillation Bottoms (ATB). ATB can be further exposed to vacuum distillation for separating vacuum gas-oils from vacuum bottoms, which are also called Vacuum Tower Bottoms (VTB). The lighter, more valuable precursor materials from the atmospheric distillation and the vacuum gas-oils from the vacuum distillation can be subjected to various kinds of hydro-treatment processes for removing impurities and to further increase the value of the lighter products. The VTB constituents are high molecular-weight aromatic and non-aromatics that require further upgrading to make fuels, chemicals or other products.
The VTB can be subjected to a thermal-cracking process whereby high temperatures and pressures are used to convert the high molecular-weight compounds into smaller molecular-weight compounds that are more valuable. Thermal cracking is typically achieved by one or more of visbreaking, delayed coking, fluid coking, or fluid catalytic-cracking. These processes all create lower molecular-weight compounds that can be separated into various valuable products by boiling-point separation and/or other processes.
At least one of the challenges of thermal cracking is to create an environment where the temperatures are high enough to cause the high-molecular weight molecules to break down while regulating the generation of unstable heavy liquids and coke, a high carbon-content solid. The generation of unstable heavy (cracked) liquids, in a process such as visbreaking, can cause fouling and coke production in the equipment and in downstream processes, which in turn can limit the generation of valuable products. Increasing the thermal severity, as in the case of a coker, generates substantial quantities of coke, which is a less valuable product.
A further challenge in upgrading heavy oil is the source of hydrogen gas for hydro-treatment processes are typically produced by one or more reformer processes. For example, the hydrogen-rich gas produced in a steam methane-reformer results in the production of greenhouse gases, such as carbon dioxide (CO2). Furthermore, current upgrading systems and processes can direct valuable carbon and hydrogen into less valuable products, such as coke, and/or a waste stream, such as a flare stack.
Upgrading of difficult to process heavy-oil is typically performed through one or more thermal processors that perform carbon-rejection processes, one or more hydrogen-addition processes, or combinations of both. Some implementations of the present disclosure relate to a process for upgrading difficult to process heavy-oil feedstocks that integrates carbon rejection and hydrogen addition processes in a manner that achieves improved yields by both increasing conversion capability and substantially reducing the hydrogen input requirements. Some implementations of the present disclosure relate to an integrated process relates that directly incorporates higher hydrogen-content hydrocarbons into thermally processed heavy-oil while simultaneously controlling Toluene Insoluble Organic Residue (TIOR) levels within the process. Implementations of the present disclosure relate to an integration of heavy-oil processes that results in heavy-oil upgrading with greatly increased liquid volumetric gain while reducing the hydrogen uptake requirement, as compared with typical carbon-rejection processes and hydrogen-addition process. As will be appreciated by one skilled in the art, the direct incorporation of high hydrogen-content hydrocarbons into the thermally processed heavy-oil is not limited to just alkylation reactions. Various other types of reactions can occur during the integrated process so that the high hydrogen content hydrocarbons are incorporated and result in volumetric gains of the liquid products while reducing the hydrogen uptake requirements.
Conventional heavy oil upgrading technologies are based on either carbon rejection or hydrogen addition. Implementations of the present disclosure directly incorporate a feedstock of intermediate hydrogen-content hydrocarbons and/or high hydrogen-content hydrocarbons with difficult to process heavy-oil feedstocks—with multiple aromatic structures and low hydrogen content—to yield intermediate hydrogen-containing products. This direct incorporation can substantially reduce or eliminate the majority of the conventional hydrogen-addition processing and the associated CO2 generation. In some implementations of the present disclosure, the carbon that would otherwise have been eliminated as CO2 is incorporated into an increased volume of the liquid hydrocarbon products, which can further reduce the overall greenhouse gas impact of the process.
Implementations of the present disclosure can result in substantially lower carbon dioxide (CO2) generation per volume of produced liquid product, as compared to known processes. In some implementations of the present disclosure, the integrated processes of the present disclosure can result in a synergistic coupling of gas to liquids and heavy-oil upgrading technologies.
Some implementations of the present disclosure relate to directly incorporating higher hydrogen-content light hydrocarbons into thermally generated, difficult to process heavy-oils produced in satellite thermal processing units, such as cokers, visbreakers and/or a hydro-visbreakers. In the case of coking, a very low hydrogen content and high carbon-content petroleum coke can be isolated and the high hydrogen content light gases can be directly incorporated with the heavy thermal liquid from the coker to produce a relatively high quality hydrocarbon stream. Beyond achieving the direct incorporation within an integrated thermal processing system and/or process, use of a coker-fractionator source of heavy-oil based feedstocks can reduce or remove the volume of fractionator-tower bottoms that are recycled back into the coker-coke drum feed. This reduced recycle volume can provide an increased volume and processing capacity of the coker-fractionator unit.
Some implementations of the present disclosure relate to a method of upgrading a heavy oil feedstock including the steps of: generating a high hydrogen-content light hydrocarbon feedstock in a thermal processor; and feeding the heavy oil feedstock and light hydrocarbon feedstock into a reaction vessel to thereby incorporate the light hydrocarbon feedstock into the heavy oil feedstock to produce a mixed effluent.
Some implementations of the present disclosure relate to a method of upgrading a heavy oil feedstock including the steps of: generating the heavy oil feedstock in a thermal processor; and feeding the heavy oil feedstock and a high hydrogen-content light hydrocarbon feedstock into a reaction vessel to thereby incorporate the light hydrocarbon feedstock into the heavy oil feedstock to produce a mixed effluent.
Some implementations of the present disclosure relate to a method of upgrading a heavy oil feedstock including the steps of: generating the heavy oil feedstock and a high hydrogen-content light hydrocarbon feedstock in a thermal processor; and feeding the heavy oil feedstock and the light hydrocarbon feedstock into a reaction vessel to thereby incorporate the light hydrocarbon feedstock into the heavy oil feedstock to produce a mixed effluent.
Some implementations of the present disclosure relate to a method of upgrading a heavy oil feedstock including a step of: feeding the heavy oil feedstock, a light hydrocarbon feedstock, and H2 into a reaction vessel to thereby incorporate at least some of the light hydrocarbon feedstock into the heavy oil feedstock to produce a mixed effluent.
Some implementations of the present disclosure relate to a system for upgrading a heavy oil feedstock. The system includes: a thermal processor; a reaction vessel; a light hydrocarbon feedstock conduit configured to feed a high hydrogen-content light hydrocarbon feedstock from the thermal processor to the reaction vessel; and a heavy oil feedstock conduit to feed the heavy oil feedstock into the reaction vessel.
Some implementations of the present disclosure relate to a system for upgrading a heavy oil feedstock. The system including: a thermal processor; a reaction vessel; a heavy oil conduit configured to transfer the heavy oil feedstock from the thermal processor to the reaction vessel; and a light hydrocarbon feedstock feed to feed a high hydrogen-content light hydrocarbon feedstock into the reaction vessel.
Some implementations of the present disclosure relate to a system for upgrading a heavy oil feedstock. The system including: a thermal processor; a reaction vessel; a heavy oil conduit configured to transfer the heavy oil feedstock from the thermal processor to the reaction vessel; and a light hydrocarbon feedstock conduit to transfer a high hydrogen-content light hydrocarbon feedstock from the thermal processor into the reaction vessel.
Some implementations of the present disclosure relate to a system for upgrading a heavy oil feedstock. The system including: a thermal processor; a reaction vessel; a heavy oil conduit configured to transfer the heavy oil feedstock from the thermal processor to the reaction vessel; a light hydrocarbon feedstock conduit to transfer a high hydrogen-content light hydrocarbon feedstock from the thermal processor into the reaction vessel; and an H2 feed to feed an H2 source into the reaction vessel.
Some implementations of the present disclosure relate to a reactor unit for upgrading a first hydrocarbon-feedstock. The reactor unit includes a first end, a second end and a sidewall that defines a plenum between the first end and the second end. The reactor unit also includes a feedstock inlet, a first gas-inlet and a first outlet. The feedstock inlet is configured to introduce a low hydrogen-content hydrogen feedstock and an anti-coking additive into the plenum proximal the first end. The first gas-inlet is configured to introduce a high hydrogen-content light hydrocarbon into the plenum at an inlet temperature of at least about 800° F. The first outlet is configured to remove a mixed effluent from the plenum proximal the second end.
Some implementations of the present disclosure relate to a system for upgrading a difficult to process heavy-oil feedstock. The system includes a reactor unit according implementations of the present disclosure; a first separator that is configured to receive and to separate the mixed effluent into a first liquid-stream and a first vapor-stream; a first hydrotreater that is configured to receive the first vapor-stream and/or a vacuum unit light product stream for increasing a hydrogen content thereof as a first hydrotreater product; a second separator that is configured to receive and to separate the first hydrotreater product into a second liquid-stream and a second vapor-stream; a third separator that is configured to receive and separate the second vapor-stream from the second separator into a third liquid-stream and a third vapor-stream; and a product fractionator that is configured to receive at least a portion of the third-liquid stream and to produce products.
Some implementations of the present disclosure relate to a system for upgrading a difficult to process heavy-oil feedstock. The system includes a reactor unit according to implementations of the present disclosure; a first separator that is configured to receive and to separate the mixed effluent into a first liquid-stream and a first vapor-stream; a second separator that is configured to receive and to separate the first vapor-stream into a second liquid-stream and a second vapor-stream; a third separator that is configured to receive and separate the first liquid-stream into a third liquid-stream and a third vapor-stream; and a second reactor unit. The second reactor unit has a first end; a second end; a sidewall that defines a plenum between the first end and the second end; an inlet that is configured to receive the third liquid-stream and/or the low hydrogen-content hydrogen feedstock; an additive inlet that is configured to introduce an anti-coking additive into the plenum proximal the first end; a first gas-inlet that is configured to introduce a high hydrogen-content light hydrocarbon into the plenum at a temperature of at least about 800° F. between the second end and the feedstock inlet, and an outlet that is configured to remove a mixed effluent from the plenum proximal the second end.
Some implementations of the present disclosure relate to a method of upgrading a difficult to process heavy-oil feedstock that includes steps of directly incorporating a first low molecular weight hydrocarbon feedstock into a thermally processed heavy-oil feedstock for producing a mixed effluent; performing at least one separating step on the mixed effluent for producing a liquid stream and a gas stream; and separating the gas stream into one or more products.
Some implementations of the present disclosure allow for upgrading of difficult to process heavy-oil feedstocks. The implementations of the present disclosure can use intermediate hydrogen content hydrocarbons and/or high hydrogen content hydrocarbons for upgrading the difficult to process heavy-oil feedstocks rather than other sources of hydrogen that produce CO2. The implementations of the present disclosure increase the hydrogen content of the feed through the direct incorporation of the high hydrogen content hydrocarbons. The present disclosure provides systems and processes to upgrade difficult to process heavy-oil feedstocks that is not carbon rejection or hydrogenation, but rather is systems and processes that directly integrate higher hydrogen content hydrocarbons into the difficult to process heavy-oil feedstocks for generating pipeline transportable products.
Some implementations of the present disclosure relate to a process that provides a sizable volumetric gain of liquid products through both direct incorporation and expansion of the product volume by the generation of smaller, less dense molecules through cracking. Many of the difficult to process heavy-oil feedstocks include metals that can impair downstream refining processes. In some implementations of the present disclosure, these metals can be utilized as a quasi-catalyst during processing. The metals can be incorporated into an ash product that can act as a catalyst within linked steps (or satellite processes) of the process. Ultimately, the metals can be isolated as a metal concentrate product. In some implementations of the present disclosure, a heavy gasoil can be generated and used as a carrier medium for moving the ash between different process steps or processes so that the ash containing stream can act as both a catalyst and a hydrogen-donor source.
Without being bound by any particular theory, implementations of the present disclosure can provide direct incorporation of one or more rich fuel gases, such as C1 through C7, into lower hydrogen content heavy oils to produce valuable intermediate hydrogen-content products. This direct incorporation can reduce the overall CO2 production that is typically associated with operating heavy-oil processes by reducing the reliance on sources of hydrogen that are associated with the production of CO2. In order to optimally utilize the relatively higher hydrogen content rich fuel gas as the hydrogen source, a heavy oil can be utilized as a feedstock. Preferably, this heavy oil would have a high resin to asphaltene ratio to inhibit the asphaltenes within the heavy oil from coking. This mitigation of the coking reactions in turn allows for increased asphaltene conversion, stability of the cracked products at elevated operating temperatures, and incorporation of higher hydrogen content feeds to the reaction system products. However, as will be appreciated by the person skilled in the art, the feedstock is not limited to feedstocks with any specific resin to asphaltene ratio. Increased asphaltene cracking can provide capping sites on the cracked hydrocarbons where the low molecular-weight, higher hydrogen content hydrocarbons can be directly incorporated into the cracked hydrocarbon-products, for example through alkylation reactions. This direct incorporation of the low molecular-weight, higher hydrogen content hydrocarbons onto these asphaltene structures can provide a significant volumetric boost to the cracked hydrocarbon-products by increasing the mass of both the carbon and hydrogen incorporated into the cracked hydrocarbon-products.
Some features, such as conduit, flow path or processing units, of the implementations of the present disclosure are optional and some of these optional features are shown in the figures with hashed lines.
Implementations of the present disclosure relate to systems and processes that produce valuable liquid hydrocarbon products from difficult to process heavy-oil feedstocks, as defined herein below. The difficult to process heavy oil feedstock can contain a sufficiently high resin to asphaltene ratio so that the resin content of the feedstock protects the asphaltene content from precipitating out of solution. The resins and/or high-boiling polar aromatics can help maintain the asphaltenes that are participating in reactions within the bulk solution during a thermal upgrading process according to implementations of the present disclosure. Minimizing asphaltene partitioning can facilitate the formation of alkylation bonding sites for the direct incorporation of lighter hydrocarbons with a medium and/or high hydrogen-content, such as lighter hydrocarbons that can be produced from the elsewhere in the upgrading facility. An example of such light hydrocarbons are the rich fuel gases produced by a coker-fractionator tower of the upgrading facility. This direct incorporation of the lighter hydrocarbons can provide an increased volume of the valuable liquid hydrocarbon products, as compared to when there is no direct incorporation. The increased volume arises from carbon and hydrogen atoms from the lighter hydrocarbons being added to the carbon chains that ultimately form part of the valuable liquid hydrocarbon product. In some implementations of the present disclosure, the lighter hydrocarbons have a higher hydrogen-content and this contributes towards generating a greater volumetric gain of the liquid products while substantially reducing or eliminating the need for hydrogen that is generated by carbon dioxide (CO2) producing processes. In some implementations of the present disclosure a recycling loop within the upgrading facility, for example a recycling loop that delivers coker-fractionator tower bottoms back upstream of the coker unit, can be decoupled (at least partially) so that the difficult to process heavy oil feedstocks are processed by the implementations of the present disclosure. The decoupling of one or more recycling loops may increase the operational capacity and operational life of various components of the upgrading facility. In some further implementations of the present disclosure, the light hydrocarbons can be supplemented with another source of hydrogen.
Some implementations of the present disclosure combine the use of anti-coking additives and other TIOR management features to facilitate a multitude of systems and processes for upgrading difficult to process heavy-oil feedstocks. This technology extends beyond the capabilities of known heavy oil upgrading technologies by one or more integrated systems and/or one or more integrated processes. For example, some implementations of the present disclosure relate to use of a coker-fractionator unit that results in increased yields, while exploiting the coker's ability to function as a carbon-rejection system. As the hydrogen content of the coke yielded from the coker is reduced to around 4 wt %, the liberated hydrogen is yielded in the liquid products and light gases. Using some implementations of the present disclosure, these light hydrocarbon gases produced from the coker-fractionator unit have a higher hydrogen-content that and can be processed with various aromatic liquids to yield intermediate hydrogen content products. Using this approach, intermediate hydrogen content products can be made while substantially reducing or eliminating the requirement to generate a hydrogen intermediate that generates associated CO2. In other implementations of the present disclosure, the process can be used with one or more products from a visbreaker-type process, where all the carbon is yielded as either a gas or liquid product. Implementations of the present disclosure can substantially reduce the production of greenhouse gases that are associated with the hydrogen addition during processing of difficult to process heavy-oil feedstocks.
Implementations of the present disclosure relate to direct incorporation of high hydrogen-content light hydrocarbons into difficult to process heavy-oil feedstocks while operating in a low hydrogen partial-pressure environment. The low hydrogen partial-pressures can provide a use for low hydrogen content streams, such as hydrotreater purges, which can reduce the need to purge gas and, thereby, reduce the energy and hydrogen that are wasted in association with typical hydrogen-processing equipment. As further described below, increasing the partial pressure of the non-pure hydrogen components of the gas can result in increasing the efficiency of the incorporation of the hydrocarbon gas into the liquid.
Implementations of the present disclosure relate to an integrated process that is capable of directly combining difficult to process heavy-oils and light gases to produce synthetic crudes and other refined liquid products. The implementations of the present disclosure can enable improved economics and facilitate transportation of the upgraded products at greatly reduced generation of greenhouse gases and at a reduced energy intensity.
Many of the difficult to process heavy-oil feedstocks include metals that can impair downstream refining processes. In some implementations of the present disclosure, these metals can be utilized as a quasi-catalyst during processing. The metals can be incorporated into an ash product that can act as a catalyst within linked steps (or satellite processes) of the process. Ultimately the metals can be isolated as a metal concentrate product. In some implementations of the present disclosure, a heavy gasoil can be generated and used as a carrier medium for moving the ash between different process steps or processes so that the ash containing stream can act as both a catalyst and a hydrogen-donor source.
Definitions
As used herein, the term “about” refers to an approximately +/−10% variation from a given value. It is to be understood that such a variation is always included in any given value provided herein, whether or not it is specifically referred to.
As used herein, the term “conduit” refers to a pipe, fluid transmission line or other mechanism for providing fluid communication between two features of the present disclosure. In some implementations of the present disclosure, use of the singular “conduit” can include multiple “conduits”. The term “conducting” may be used interchangeably with the terms “feeding” or “flowing” and these terms refer to the movement of a fluid, with or without entrained solids, through a conduit.
As used herein, the term “downstream” refers to a position or component within a system, apparatus, unit or a step within a process that is after a prior position, component or step.
As used herein, the term “difficult to process heavy-oil feedstock” can be used interchangeably with “difficult to process heavy oil” and both terms refer to hydrocarbons with multiple aromatic structures and low hydrogen content, including but not limited to: petroleum crude oil; heavy cycle oils; shale oils; heavy oil; bitumen; high-boiling point fractions and solid fractions that are separated from heavy oil or thermally-generated components from heavy oil upgrading; vacuum-tower bottoms (VTB); coker-fractionator bottoms; coker heavy-gasoil; mid to high nC7 asphaltenes; low hydrogen-content hydrocarbons; aromatic hydrocarbons; mid to high polar hydrocarbons; coker gas oil such as HVGO; visbreaker bottoms; hydro-visbreaker bottoms, a mixture of components like a diluent and a heavy oil, where the diluent can be a C5C6 type diluent that is mixed with Athabasca Bitumen (Western Canadian Select), the diluent can also be other light hydrocarbons that are used in crude or bitumen solvent extraction processes and that are mixed with a difficult to process heavy oil; other products of thermal processing of heavy oil; or, combinations thereof.
As used herein, the term “high hydrogen-content” refers to hydrocarbons that have a wt % of hydrogen that is higher than an intermediate hydrogen-content range. Some non-limiting examples of high hydrogen-content hydrocarbons include, but are not limited to: coker naptha; visbreaker naptha; and, combinations thereof. One skilled in the art will also appreciate that these terms regarding the hydrogen content can also be used as a more general reference between the different streams and sources of hydrocarbons described herein.
As used herein, the term “intermediate hydrogen-content” refers to hydrocarbons with a weight percent (wt %) of hydrogen between about 11.5 wt % and about 13 wt %. One skilled in the art will also appreciate that these terms regarding the hydrogen content can also be used as a more general reference between the different streams and sources of hydrocarbons described herein.
As used herein, the term “low hydrogen-content” refers to hydrocarbons that have a wt % of hydrogen that is lower than the intermediate hydrogen-content range. One skilled in the art will also appreciate that these terms regarding the hydrogen content can also be used as a more general reference between the different streams and sources of hydrocarbons described herein.
As used herein, the term “rich fuel gas” refers to low molecular-weight hydrocarbons, which can also be referred to as light hydrocarbon gases, that contain hydrogen and they include, but are not limited to: C1 gas; C2 gas; C3 gas; C4 gas; C5 gas; C6 gas; refinery fuel gas; light hydrocarbons from a second stage of a system 702 (as described further herein below); fuel reformer hydrogen gas; FCCU fuel gas; FCCU C3, C4, C5; gas field products C1, C2, C3, C4, C5, C6; coker product light ends C1, C2, C3, C4, C5, C6; visbreaker product light ends C1, C2, C3, C4, C5, C6; hydrotreater purge gas; hydrocracker purge gas; hydrogen product unit raw gas; or, combinations thereof. In comparison to the difficult to process heavy-oil feedstocks discussed herein, the rich fuel gases can have a high hydrogen content.
As used herein, the term “upstream” refers to a position or component within a system, apparatus, unit or a step within a process that is before a subsequent position, component or step.
Coker-Fractionator System
As shown in the non-limited example of
The vacuum distillation tower 14 applies a vacuum pressure to the atmospheric bottoms for extracting the light vacuum gas oils, heavy vacuum gas oils from the vacuum tower bottoms. The light vacuum gas oils can be conducted by a conduit 110 to combine, or not, with the atmospheric light gas oils for further processing. The heavier vacuum gas oils can be conducted by a conduit 112 or 114 from the vacuum distillation tower 14, also for further processing. The vacuum tower bottoms are conducted away from the bottom of the vacuum distillation tower 14 by a conduit 116A.
In the example upgrading system 10 shown in
Within the coker drum 21, the coker feedstock can be heated and pressurized to produce a coker product through a thermal-cracking process. The coker product is made up of cracked hydrocarbon vapor and entrained solid coke-particles, the cracked hydrocarbon vapor can also be referred to as a cracked hydrocarbon vapors product or a coker drum effluent. The cracked hydrocarbon vapor can include a wide range of constituents including non-hydrocarbons and hydrocarbons. The non-hydrocarbons constituents can include, but are not limited to: hydrogen (H2) and hydrogen sulfide (H2S). The hydrocarbons constituent within the cracked hydrocarbon vapor can include, but are not limited to: methane (CH4), C2 to C4 hydrocarbons, a naphtha fraction, a kero fraction, and a gas oil fraction. The boiling point of the hydrocarbon constituents of the cracked hydrocarbon vapor can be in excess of 1000° F.
The solid coke-particles can also be referred to as coke or petroleum coke. The solid coke-particles include micro-carbon content that reflects the amount of heavy hydrocarbons with a high coking tendency. There are two types of micro-carbon. One type is referred to as distillable micro-carbon, which is generated by the hydrocarbons that are vaporized at the coker-fractionator unit's normal operating temperatures. The other type of micro-carbon is referred to as non-distillable micro-carbon, which is generated either by the hydrocarbons that cannot be distilled due to a high boiling-temperature, the presence of a multi-ringed structure, or the non-distillable micro-carbon can also be the coke fine itself. The non-distillable micro-carbon can end up in the fractionator tower 22 hydrocarbon products, as described further below, due to carry-over or entrainment within vapor streams within the coker-fractionator unit 200.
The coker product exits the coker drum 21 by the CVL 217, which conducts the coker product into the fractionator tower 22. In some implementations of the present disclosure, the CVL 217 can be between 500 and 2000 feet long (one foot is equal to about 0.305 meters). In some implementations of the present disclosure, substantially most of the solid coke-particles remain within the coker drum 21 but at least a portion of the solid coke-particles can become entrained within the stream of cracked hydrocarbon vapor and the entrained particles can be conducted by the CVL 217. In some examples of a coker-fractionator unit 200, the contents of the CVL 217 have a temperature of about 900° F. and a pressure of about 40 pounds per square inch gauge (psig, which is substantially equal to about 377 kilo-Pascals).
Solid coke can be removed from the coke drum 21 by known methods, which are collectively represented by line 230.
Within the fractionator tower 22 the coker product is separated into a top vapor product that is conducted by a conduit 218 that contains coke gas and rich fuel gases. The coker product is also boiling-point separated into further vapor products that are conducted away from the fractionator tower 22 by conduits 221. For example, the further vapor products include light coker naphtha (within a conduit 222), heavy naphtha (within a conduit 224), coker kerosene (within a conduit 226) and coker gas oil (within a conduit 228).
The fractionator tower bottoms have a high sulfur, nitrogen and oxygen content and, therefore, the fractionator tower bottoms are very polar. Additionally, the fractionator tower bottoms are very low in hydrogen content and they include many multi-ring aromatic structures. Due to these chemical properties, the fractionator tower bottoms can be difficult to process further. Typically, the fractionator tower bottoms are recycled back upstream of the coker drum 21 to combine with the coker feedstock within the conduit 116A via a recycle conduit 232. The recycle conduit 232 can continuously introduce a desired volume, over a specified time, of the fractionator tower bottoms into the coker drum 21 so that the recycled fractionator tower bottoms are continuously recycled until they are coked within the coker drum 21. This desired volume of recycled fractionator tower bottoms occupies a given volume of the coker drum 21, which necessarily reduces the volume of new coker feedstock that can be introduced into the coker drum 21 over a specific time.
As will be appreciated by one skilled in the art, the flow rate within the recycle conduit 232 can set the temperature cut-point for the further vapor products within the conduits 221, which can influence the quality of the further vapor products that are sent to downstream hydrotreaters, or other processing units, for further processing.
In some implementations of the present disclosure, the reaction unit 24 can perform a thermal upgrading process, which may also be referred to as a thermal cracking process. For example, the unit 24 can be a slurry-phase hydrocracking reaction vessel (also referred to as a SHC unit) within which a slurry-phase hydrocracking upgrading process (a SHC process) can occur. In other implementations of the present disclosure the unit 24 can be an upgrading system 700 or an upgrading system 702, each of which include at least one integrated thermal processing will be described further herein below. In other implementations of the present disclosure the unit 24 can include a SHC unit and one or both of the system 700 and the system 702.
Implementations of the present disclosure can provide an economic way to upgrade low-value feedstocks, such as difficult to process heavy oil feedstocks. Without being bound by any particular theory, the coker-fractionator unit 200A can have at least the following advantages over the coker-fractionator unit 200: protection and longer operational life for the coker unit and downstream fixed-bed catalysts; coker-yield increases through a reduced pressure within the coker drum 21 and directing the fractionator-tower bottoms for further processing within the unit 24. In some implementations of the present disclosure, the unit 24 permits further processing of the fractionator tower bottoms, as an example of a difficult to process heavy-oil feedstock, rather than just recycling via the conduit 232 until the fractionator tower bottoms are converted to coke and gas. By diverting some or all of the content of the conduit 232 to the unit 24, the coker drum 21 can have increased volumetric capacity, which can also increase the coker yields. Furthermore, the rich fuel gases can serve as a less expensive source of hydrogen than pure hydrogen. On a BTU basis, the rich fuel gases are typically sold at a steep discount to other oil products. Through implementations of the present disclosure, the rich fuel gases can be directly incorporated into the difficult to process heavy-oil feedstocks to generate products that can be sold as a synthetic crude oil or a refined liquid product.
SHC Process
In some implementations of the present disclosure, the SHC process that occurs within the reaction unit 24 is a thermal upgrading process that includes the use of anti-coking additives, a hydrogen-gas stream and a pressurized, high temperature vessel used to upgrade heavy-oil feedstocks within a slurry phase. The SHC process uses polar aromatic compounds for slowing the self-association of toluene insoluble organic residues (TIOR) during the thermal upgrading thereby causing an increased upgrading potential of a given feedstock, while yielding less or no coke. During the thermal upgrading of the feedstocks in the SHC process, the asphaltene exists in association with resins, which are smaller, polar-aromatic structures and other higher hydrogen-content hydrocarbon structures. The asphaltenes crack at slower rates relative to surrounding hydrocarbon structures, which maintain the asphaltenes in a suspension. The asphaltenes contain the highest concentration of the oxygen, nitrogen and sulphur polar species relative to the other hydrocarbon species. As the heavy oil is upgraded, the polar species in the asphaltenes can become concentrated due to the thermal cracking of the asphaltenes. The combined effect of the more rapid conversion of the supporting resin hydrocarbons relative to the asphaltenes and the concentration of the more polar asphaltene species result in their tendency to self-associate and form mesophase coke and TIOR. Factors that favour relative increased rates of hydrogenation of the resins, such as active catalyst systems, result in increased mesophase coke generation. The addition of the polar aromatic resin type structures into the feedstock limits the asphaltenes self-association and generation of mesophase coke, while these compounds undergo thermal conversion, by performing a function similar to the original resins within the feed.
It is also known that the use of polar aromatic resins can reduce the TIOR and the conversion of TIOR to ash. Ash is a combination of an iron sulfide (FeS) anti-coking additive; metals laid down from conversion of feedstocks; and any solids from the feedstock such as silt. The higher polarity TIOR compounds are associated with the FeS anti-coking additives. As the TIOR concentration increases, the individual ash and TIOR particulates in the slurry suspension associate together creating larger, denser particles that settle out causing coke laydown in the reaction and fractionation systems.
During the SHC process, the anti-coking additives and the feedstock input for the SHC process are premixed, heated and added into the SHC unit to form a slurry-phase. Within the slurry-phase, the anti-coking additives inhibit the formation of coke. The mixture of heavy oil input and anti-coking additives are introduced by input feed nozzles located at the bottom of an SHC vessel within the SHC unit. In some SHC units, the hydrogen-gas stream is heated to between about 842° F. and about 1112° F. and introduced by gas feed nozzles with a velocity of at least 390 ft/sec located above the input feed nozzles within the SHC vessel. The liquid feed and some gas is introduced at the bottom of the SHC vessel at between about 572° F. and about 806° F. and above a velocity of about 82 ft/sec. The thermal and kinetic energy of these two streams provide the energy for the cracking and hydrogenation reactions, mixing, and vaporization of the light hydrocarbons generated.
Some SHC configurations have also demonstrated the ability to upgrade solid carbonaceous materials into liquid products. This ability to upgrade coal and petroleum coke differentiates the SHC configurations from the other thermal upgrading processes.
Integrated Thermal Process (ITP)
Briefly, the gas-contacting system 904 system includes components that are useful for processing harder-to-vapourize, high viscosity, high-boiling feedstocks. The gas-contacting system 904 physically prepares the hydrocarbons exiting the gas inlet 902, the contents entering through the inlet 902, and the most concentrated TIOR-Ash segment of the contents within the reactor unit 30. The gas-contacting system 904 provides the source of the high temperature gas with temperatures typically in excess of about 800° F., about 900° F., about 1000° F. or higher. In some implementations of the present disclosure, the gas temperature at the inlet 900 can be more than 200° F. hotter than the bulk reactor temperature in the reactor unit 30. As the molecular weight of the high hydrogen content gas is introduced through gas inlet 902 increases, the amount of energy both in terms of enthalpy and kinetic energy at the discharge of the gas inlet 902 is increased. At any given velocity through the gas inlet 902, the gas jet penetration also increases energy transferred to the reactor contents in the proximity of the gas-contacting system 904. Improving the efficiency in this energy-transfer process can reduce the partitioning which reduces the TIOR yield, increases conversion, and decreases the gas yield. Within the reactor unit 30, the hydrocarbon vapour contact time is typically in the range of about 1 to 2 minutes. However, the contact time for the reactor contents to quench the high temperature gas jets exiting 902 is in the order of milliseconds. As the temperature is increased within the reactor unit 30, the contact time is reduced for a given feedstock conversion. The maximization of this contact temperature and the minimization of this contact time with the rapidly quenched gas jets can result in the maximization of the olefinic reactions, which can impact subsequent availability for further reaction pathways. The configuration of reactor 30 and the operating conditions are set-up to segregate and position the TIOR-Ash complex in contact with the fluids entering the reactor unit 30 by the gas inlet 902, thereby maximizing the energy intensity at the point of the maximum concentration of TIOR and Ash in the reactor unit 30. Exposure to this maximized energy intensity can be followed by an immediate quenching to the bulk reactor temperature.
In some implementations of the present disclosure, the gas-contacting system 904 can provide one or more of the following aspects to facilitate the production of ITP products from a difficult to process heavy-oil feedstock:
a. the high intensity energy associated with introducing the hydrocarbon by a gas jet;
b. the physical-contact parameters (such as mass, velocity, geometry, temperature and others) that create a situation analogous to the generation of small feed droplets within a fluid catalytic coker (FCC) riser;
c. a physical proximity of the TIOR and Ash to interact with: the gas inlet 902; the TIOR-Ash suspension at the bottom of the reactor unit 30 due to a combination of “fluidization” and particle size control; and a polar-aromatic control system. The polar-aromatic control system allows the inlet 904 to position the ash for contacting with the hydrocarbon vapor and subsequent withdrawal. In some respects this is analogous to reducing an FCCU feed viscosity to promote the generation of smaller droplets within the FCCU contacting system; and
d. a rapid quenching of the high temperature jets by the bulk solution in the reactor unit 30. The gas-contacting system 904 can become more effective as lower hydrogen content gas is introduced at the inlets 904.
Without being bound by any particular theory, the heat supplied by the high density gas jets at the inlets 902 will be at temperatures far above previous commercial thermal-upgrading operations. The reactor unit 30 can include a quench that acts to maximize the jet contacting temperature and allow the bulk solution within the reactor unit 30 to operate at a lower temperature, which can be beneficial in some configurations, such as a satellite processing unit.
The inlet 900 can also provide control over the introduction of polar aromatic oils (donor-solvents) and FeS anti-coking additives, which in turn can influence the particle size of the Ash and facilitate the direct incorporation of light hydrocarbons by alkylation reactions. In some implementations of the present disclosure, the reactor unit 30 can also include one or more densitometers 906 that are configure to monitor the density of the mixed, three phase contents of the reactor unit 30 to allow determinations of Ash and TIOR content. In some implementations of the present disclosure at least two densitometers 906 are positioned within the reactor unit 30 between the first end 35A and the second end 35B. In some implementations of the present disclosure at least one densitometer 906 is positioned within the reactor unit 30 at or proximal to the gas-inlet. In some implementations of the present disclosure at least one densitometer 906 is positioned within the reactor unit 30 between the gas-inlet and the second end. In some implementations of the present disclosure at least one densitometer 906 is positioned within the reactor unit 30 proximal to the first end 35A within one third (⅓) to two thirds (⅔) of the distance between the first end 35A and the second end 35B. By monitoring differences in the concentration of the high density TIOR and Ash within the reactor, the TIOR inventory can be directly monitored. This methodology takes advantage of the Ash gradient set-up over the length of the plenum 39 that results from the difference in settling velocities at different TIOR/Ash ratios. The concentration of Ash and TIOR at the point of the gas contactor can be optimized via the asphaltene accumulation balancing operating parameters or by solvency adjustments facilitated by adding more or less high solvency feeds into the bottom of the reactor below the gas inlet 902, or adjusting the quantity of higher hydrogen content, low solvency feedstock into the top half of the reactor unit 30. The temperature and energy input through the gas-contacting system 904 can be maximized by using a gas or liquid quench into the top of the reactor. While elevated ash content in the reactor unit 30 acts to minimize foaming in the reactor unit 30, an anti-foam agent can be injected into the top of reactor as a supplemental method for controlling foaming in the reactor unit 30. As will be appreciated by one skilled in the art, the anti-foam agent can be injected into the reactor unit 30 by one or more inlets, as can other feeds that can be desirable to include within the plenum 39 of the reactor unit 30.
In some implementations of the present disclosure, the environment within the reactor unit 30 can cause very high conversion, which in turn can cause a new issue of preventing the additive and other ash from building up within the reactor unit 30 process. The ash inventory within the reactor unit 30 can be maintained at elevated levels because that can be favourable for the desired chemical reactions and for minimizing foaming in the system. In some implementations of the present disclosure, the average concentration of ash within the reactor unit 30 is at least 15 wt % of the total contents of the reactor unit 30. In other implementations of the present disclosure, the average ash concentrations within the reactor unit 30 is greater than 17 wt %, greater than 19 wt % or greater than 21 wt % of the total contents of the reactor unit 30. At these higher ash concentrations, the reactor unit 30B can display a fluid-bed circulation profile that can enhance TIOR management with ash circulating down an annular area near the reactor unit's 30 side walls.
In some implementations of the present disclosure, the reactor unit 30 may include a recycle gas loop of one or more conduits that can conduct gaseous products within the plenum back to the gas-contactor system 904 in a recycle gas loop. In some implementations of the present disclosure, the recycle gas loop can include a slip stream for removing impurities from the recycled gas. Some non-limiting examples of impurities within the recycle gas include, but are not limited to H2S, H20 and NH3.
Single Stage ITP System
The ITP process can be implemented with a number of different system configurations and with one or more reactor units 30 similar to that shown in
The lowest hydrogen content gas should contact the most receptive alkylation bonding sites within the reactor unit 30B at the gas contactor 904. Both of the gas and liquid feedstocks are introduced at locations with that objective. The operating conditions are set-up to isolate the most receptive heavy aromatics and transport them to the gas contactor 904 to interact. Operating conditions are set-up to maximize direct incorporation of the light high hydrogen content hydrocarbons into the low hydrogen content hydrocarbon components at the gas contactor 904 and to provide sufficient energy to reduce and/or negate partitioning of the polar species.
Some implementations of the present disclosure relate to the reactor unit 30B receiving a slurry-feed mixture of the difficult to process heavy-oil feedstock and from about 0.01-4.0% by weight (based on fresh feedstock) of coke-inhibiting additive particles move upwardly from a high intensity mixing zone through a confined vertical hydrocracking zone within the reactor unit 30B. The reactor unit 30B can be maintained at a temperature of between about 660° F. (about 350° C.) and about 1150° F. (about 600° C.) at a pressure of about 3.5 mega Pascals (MPa) to about 24 MPa. In some implementations of the present disclosure, the reactor unit 30B can have a space velocity of up to 4 volumes of hydrocarbon oil per hour per volume of hydrocracking zone capacity (LHSV). Within the reactor unit 30B, the gas-contacting system 904 can include an arrangement of the gas input nozzles 902 that introduce the hydrogen containing gas with sufficient thermal energy and kinetic energy to create an environment that will break apart TIOR and facilitate the direct incorporation of the hydrogen containing gas onto the low hydrogen hydrocarbon feedstock. These gas input nozzles 902 are part of the gas contactor 904. The difficult to process heavy-oil feedstock, anti-coking additive, polar aromatics are supplemented with sufficient high hydrogen content gas to enter through inlet 900 within the reactor unit 30B to optimally distribute the TIOR-Ash complex for interaction within the gas-contacting system 304 and prevent deposition in the bottom of the reactor unit 30B.
Under these parameters, the contents of a three-phase reaction system, including the products of the conversion of the various feedstocks, recycle gas and ash exit the reactor unit 30B as a mixed effluent from the top of the reactor unit 30B by a conduit 328. In the case of very high feedstock conversion, conduits 304, 328, can contain vapour products only with an ash stream exiting the bottom of the reactor through conduit 330 (and/or conduits 390 and 392 from reactor unit 30A, as shown in
The reactor unit 30B can crack difficult to process heavy-oil feedstocks. The difficult to process heavy oil feedstocks contain various amounts of asphaltenes. As will be understood by one skilled in the art, asphaltenes are high molecular-weight compounds that contain heteroatoms, which impart polarity. Asphaltenes also contain aromatic structures and they can be highly unsaturated. Asphaltenes are also known to be surrounded by a layer of resins made up of polar aromatic structures. The resins are a mixture of lower molecular-weight class of compounds that have many of the same chemical features as the asphaltenes. The resin can stabilize the asphaltenes in colloidal suspensions. In the absence of the resin, the asphaltenes can self-associate, or flocculate to form larger molecules which can precipitate out of solution. This is the first step in coking. The difficult to process heavy oil feedstocks also have a lower ratio of resin to asphaltenes. One non-limiting example of a difficult to process heavy oil feedstock is Visbreaker bottoms derived from a Mene Mota VTB visbreaker feed, which has a resin to asphaltene ratio of about 0.56:1.
The reactor unit 30B can operate at a higher temperature and lower hydrogen partial-pressure than typical hydrocracking processes and systems. Without being bound by any particular theory, a very short contact, higher temperature reaction environment can provide an improved balance between the thermal asphaltene cracking and the cracking of the resin. A lower hydrogen partial-pressure can also result in benefits in hydrogen management. Although the ITP process can be carried out in a variety of known reactors with either up or down flow, the process is particularly well suited to a tubular vessel through which the mixture of difficult to process heavy-oil feedstock, the additive particles and a hydrogen-containing gas move upwardly due to the high mixing environment at the base of the reactor unit 30B caused by the gas contactor 904 and the auto-cooling effect of the vapourization of the lower molecular weight cracked products.
A variety of additive particles can be used in the reactor unit 30B, provided that the additive particles survive the operating temperatures and pressures of the ITP process and remain effective as part of any recycle loops. Particularly useful additive particles include FeS particles with a particle size of less than about 45 microns (μm) and with a major portion, i.e. at least 50% by weight, preferably having particle sizes of less than 10 μm. The FeS particles can be mixed with the difficult to process heavy-oil feedstock and enter into the reactor unit 30B. A portion of the heavy hydrocarbon oil product is used to form the recycle stream of the present disclosure. The particle size of the FeS introduced into the process becomes smaller and more active with time in the process. The increased activity is due to the inclusion of materials contained in the feed such as vanadium (V) and nickel (Ni) that become an integral part of the ash. Even materials such as fine sand contained in feedstocks such as some mined bitumen become active components in the ITP ash mix.
Upstream of the reactor unit 30B, the difficult to process heavy-oil feedstock is conducted from its source to a gas heater 35E via a conduit 300 for heating to temperatures between about 600° F. and 800° F. The heated feedstock is conducted to the reactor unit 30B via a conduit 374 so that the heated feedstock enters at or near the bottom of the reactor unit 30B and is proximal to the gas contactor 904. A conduit 328 conducts a mixed effluent from the top of the reactor unit 30B to a high temperature, high pressure separator 31C operating at typically between 600° F. and 800° F. The separator 31C separates the mixed effluent into a liquid and ash stream 332 and a vapor stream 342. The conduit 332 conducts the liquid and ash stream to a low temperature, low pressure separator 32B. The separator 32B can operate over a variety of temperatures and pressures to separate the liquid and ash stream into a further liquid and ash stream and a further vapor stream. The further liquid and ash stream is conducted by a conduit 334 to an optional high boiling point, sour fractionator 36. If the fractionator 36 is not present, then the liquid and ash stream can be conducted by a conduit 338 and be recycled back into the reactor unit 30B by conduit 300, or conduit 338 which can communicate with a conduit 396 to allow the stream to be communicated upstream or conduit 338 can communicate with a conduit 366 where unconverted material can be recovered and the Ash can be recovered, as discussed further below. The further vapor stream is conducted by a conduit 388 to communicate with a conduit 386 and/or a conduit 394, as discussed further below.
The vapor stream from the separator 31C is conducted by a conduit 342 to an optional first hydrotreater vessel 33A. The first hydrotreater vessel 33A can also receive a vapor stream from the high boiling point, sour fractionator 36 (if present) by a conduit 344. The optional first hydrotreater vessel 33A can also receive an optional stream of high-purity hydrogen via a conduit 376. In some implementations of the present disclosure, the optional stream of hydrogen can come from a steam-methane reformer.
The vapor stream from the separator 34 can include light gas and naphtha up to a full range of materials, depending on the temperature of the separator 34. The vapor stream can be conducted to a low temperature, high pressure separator 32C by a conduit 348. The separator 32C can produce a liquid product stream and a vapor stream and the product stream can be conducted by the conduit 360 to a product fractionator 37 that separates the product stream into further valuable product streams, for example by boiling point separation or other applicable methods. The vapor stream from the separator 32C can be conducted by a conduit 351 to a gas heater 35D and the heater 35E. The gas heater 35D can heat the vapor stream from the separator 32C to a temperature between about 800° F. and 1400° F. and this heated vapor stream can be conducted by a conduit 325 to enter the reactor unit 30B. In some implementations of the present disclosure, the product fractionator 37 can also produce a stream of high hydrogen content light hydrocarbon feed that is conducted back into the reactor unit 30B via a conduit 361, which can pass through one or more heaters before entering into the plenum 39 of the reactor unit 30B.
In some implementations of the present disclosure, the liquid stream is conducted by a conduit 352 to a product finishing hydrotreater system 602 that includes a second hydrotreater vessel 33B and a low temperature, high pressure separator 32D. The second hydrotreater 33B in turn produces a liquid stream and a vapor stream. The second hydrotreater vessel 33B can also receive an optional stream of high-purity hydrogen via a conduit 354. In some implementations of the present disclosure, the optional stream of hydrogen can come from a steam-methane reformer. The liquid stream from the second hydrotreater vessel 33B is conducted by a conduit 356 to the separator 32D. The vapor stream from the second hydrotreater vessel 33B can be communicated with the contents of the conduit 388. The separator 32D can produce a product stream and a vapor stream and the product stream is conducted by a conduit 358 to communicate with a conduit 360. The vapor stream from the separator 32D can be communicated with the contents of the conduit 388. In other implementations of the present disclosure, the vapor stream from the separator 34 is not conducted to the product finishing hydrotreater system 602, rather the vapor stream is communicated with the contents of the conduit 348.
This vapor stream can bypass loop 602, if present, and head directly to the separator 32C.
In some implementations of the present disclosure, a stream of high hydrogen-content materials can be conducted by a conduit 386 to communicate with the contents of either or both of the conduit 351 and the conduit 328. The conduits 388 contain the recycled gases from various separators in the process which communicate with conduit 386. The conduit 351 contains the recycle gas from separator 32C. The conduit 328 contains the mixed effluent from the ITP reactor unit 30B. The conduit 351 enters either the heater 35D or the heater 35E via conduit 372. What enters the heater 35D are the recycle gas from the ITP process and the high hydrogen content material. Some examples of high hydrogen-content materials includes: gas field products such as C1, C2, C3, C4, C5, C6 and the like; FCCU derived fuel-gas, such as H2, C1, C2, C2 olefins (C2o), C3, C3 olefins (C3o), C4s, C4 olefins (C4o); coker derived fuel-gas, such as H2,C1,C2,C2o,C3,C3o,C4s,C4o; visbreaker derived fuel-gas, such as H2,C1,C2,C2o,C3,C3o,C4s,C4o; purge gases from hydrotreaters, such as H2,C1,C2,C3; light hydrocarbons from downstream unit separators, the contents of conduit 388 and combinations thereof. The high hydrogen content material will exit the heater 35D via conduit 325 and enter the reactor unit 30B. Conduit 325 directly supplies inlet 902 detailed in
In some implementations of the present disclosure, a stream of intermediate hydrogen-content materials can be conducted from a source to enter the reactor unit 30B by a conduit 384. The conduit 384 can inject intermediate hydrogen-content materials that provide quenching, facilitate greater TIOR management, and supplies additional carbon and hydrogen to the reactions within the reactor unit 30B. For example, the intermediate hydrogen-content material could be a paraffinic crude VTB with a hydrogen content of about 12 wt % and could contain light hydrocarbons.
In some implementations of the present disclosure, a stream of high hydrogen-content materials can be conducted from a source to enter the reactor unit 30B by a conduit 382. The high hydrogen-content materials can be one or more of coker naphtha, visbreaker naphtha, flashed low boiling diluent from diluted bitumen or combinations thereof. The high hydrogen-content materials can enter at or above the gas-contacting system 904 of the reactor unit 30B. In some implementations of the present disclosure, the high hydrogen content materials can act as a quench within the reactor unit 30B. The ITP reactor unit 30B utilizes the quench to reduce the temperature within the reactor unit 30B. This reduction in reactor temperature allows the gas-contacting system 904 to inject more gas or gas at a higher temperature facilitating heavy aromatic conversion and direct incorporation of the high hydrogen hydrocarbon content into the feedstock.
The reactor unit 30B also produces a liquid stream, which can also be referred to as a reactor drag-stream, that can be conducted by a conduit 330 to one or more of satellite processing units 604 by a conduit 370 and/or to a metal reclamation unit 606 by a conduit 368. The conduit 330 can have an outlet within the reactor unit 30B that is positioned above the gas contactor 904, for example between about 1 and 5 feet above the gas contactor 904, or more about 3 feet above the gas contactor 904. In some implementations of the present disclosure, the outlet for conduit 330 can be positioned within the bottom half (½) or bottom (⅓) or bottom quarter (¼) of the height of the reactor unit 30B.
The satellite processing units 604 can further process the liquid stream from conduit 330. For example, the one or more satellite processing units 604 can be a coker-fractionator unit, a visbreaker unit or a hydro-visbreaker unit. The metal reclamation unit 606 can isolate metals, such as Nickel (Ni) and/or Vanadium (V) in the liquid stream from the reactor unit 30B. Additionally, the conduit 330 from the reactor unit 30B can contain TIOR materials and Ash, so the conduit 370 can provide these materials to the one or more satellite processing units 604. For example, excessive amounts of TIOR materials can be sent to a satellite coker-fractionator unit for further high temperature carbon rejection processing. In other examples, high conversion/low TIOR polar aromatic materials and ash can be sent to one or more satellite processing units 604 as a hydrogen donor for increasing conversion within those satellite processes. Metals such as V, Ni, Iron (Fe), Titanium (Ti), Chromium (Cr), Manganese (Mn), Magnesium (Mg), Molybdenum (Mo), Strontium (Sr), Cobalt (Co), Zinc (Zn), or combinations thereof can be isolated via a clarifier or other known approaches. These metals are transported out with gas oil. This gas oil/ash mixture is transported to a low pressure clarifier. In the low pressure clarifier, the highly viscous gas oil is readily separated from the ash. The ash falls to the bottom where the remaining hydrocarbon can be burned. What remains is an oxide isolate of these various metals that is hydrocarbon-free.
In the implementations that have a high boiling point, sour fractionator unit 36, a liquid stream is generated therein that is conducted by either or both of a conduit 341 and a conduit 362. The conduit 341 conducts at least a portion of the liquid stream from fractionator unit 36 to communicate with a conduit 336 and/or a conduit 378. The conduit 336 conducts its contents to a source 38 of anti-coking additive, for example the FeS-based additive as described herein above. In other implementations of the present disclosure, the source 38 of additive can include raw anti-coking additive and/or a polar-aromatic carrier material. The liquid stream within the conduit 336 can be a carrier for conducting the anti-coking additive into a conduit 378 and/or a conduit 380. The conduit 378 conducts its contents to join the primary feedstock within the conduit 300. The conduit 380 communicates its contents to join a conduit 364. The conduit 364 conducts its contents, which can include polar aromatic compounds and Ash to one or more satellite processing units 604. The conduit 362 conducts at least a portion of the liquid stream from the fractionator unit 36 to communicate with the conduit 364 and/or a conduit 366. The conduit 366 conducts its contents to the metal reclamation unit 606. The conduit 364 conducts its contents to conduit 396 which continue on to one or more of the satellite processing units 608.
Two Stage ITP System
The first unit 30A can generate a mixed effluent that is conducted by a conduit 304 to a first high temperature, high pressure separator 31A. The first separator 31A produces a liquid stream and a vapor stream. The liquid stream can be conducted by a conduit 308 to a second high-temperature, high pressure separator 31B. The second separator 31B produces a further liquid stream and a further vapor stream. The further liquid stream can be conducted by a conduit 318 to communicate with the contents of a conduit 322, which will be discussed further below. The further vapor stream can be conducted by a conduit 310 to a first low temperature, high pressure separator 32A. The separator 32A also receives the vapor stream from the first separator 31A by a conduit 306. The separator 32A produces a liquid stream and a vapor stream. The vapor stream can be conducted by a conduit 312 into a conduit 316, as will be discussed further below. In some implementations of the present disclosure, the vapor content of conduit 316 can include one or more gases with medium hydrogen-content and/or high hydrogen-content. The liquid stream from the separator 32A can be conducted by a conduit 314 to communicate with the contents of the conduit 328 of the loop 600.
In some implementations of the present disclosure, the second unit 30B is configured to receive ash from the separator 31B via conduit 318 and optionally from the first reactor drag stream within the conduit 390. The ash within conduit 318 and within conduit 390 can be produced within the systems of the present disclosure with a smaller average particle size and so they are generally more active in hydrogen transfer reactions than the anti-coking additive systems (such as the FeS additive system). In addition to the smaller average particle size of the ash within the conduits 318 and 390, this ash can have a lower TIOR:ash ratio because this ash has already been at least partially processed by the separator. In this context, the contents of the conduit 392 could be rich in an easier to separate material, such as sand or silt, and this material that could be suitable for disposal in a coker as coke.
The first unit 30A can also produce a first-reactor drag stream that can be conducted by a conduit 390 to communicate with the contents of the conduit 392 and/or with a conduit 322. The conduit 322 conducts its contents into the second unit 30B. The first-reactor drag stream can provide inventory balancing of the TIOR materials and the Ash. The first-reactor drag stream can also provide a mechanism by which the TIOR materials are transported within a medium of polar aromatic oil. The drag stream can be sent via conduit 390 through conduit 392 and into conduit 370 to reach the external processing units, 604, where the stream can be utilized to enhance the conversion of the previously mentioned external processing units.
The conduit 312 can communicate with the contents of a conduit 316 that can conduct its contents to a feed heater 35B, which are heated and conducted by a conduit 320 to enter the first unit 30A. The conduit 316 can also communicate with a conduit 371 that conducts its contents to communicate with the primary feedstock in the conduit 300.
Other differences between the system 700 and the system 702 include that in the system 702: the contents of the conduit 350 can be communicated with the contents of the conduit 316; the contents of conduit 394 can be communicated with the contents of the conduit 312; and, the contents of the conduit 364 can be communicated with the primary feedstock within the conduit 300.
In some implementations of the present disclosure, the system 702 does not include the heater 35D but at least some of the contents of the conduit 350 can be conducted to a gas heater 35A by a conduit 324 and then the heated contents can be conducted into the second unit 30B.
In some implementations of the present disclosure, the system 702 does not include the heater 35E but at least some of the contents of the conduit 350 can be conducted to a conduit 322 by conduits 324 and 327 and then conducted into the second unit 30B.
In some implementations of the present disclosure, the conduit 384 can communicate intermediate hydrogen-content materials into the first unit 30A.
In some implementations of the present disclosure, the conduit 382 can communicate the higher hydrogen-content materials into the first unit 30A. The conduit 382 can provide higher hydrogen-content hydrocarbons into the system 702 as either liquids or vapours.
The conduit 394 can provide higher hydrogen-content hydrocarbons into the system 702. The higher hydrogen-content hydrocarbons will be injected into the first unit 30A through the 902. The contents of the conduit 394 can be one or more of the same constituents of the higher hydrogen-content materials within the conduit 382.
In some implementations of the present disclosure, a conduit 386 can communicate higher hydrogen-content materials with the contents of the conduit 304.
In some implementations of the present disclosure, a conduit 386 can communicate distillates and lower boiling liquids materials with the contents of the conduit 304 for the purpose quenching the 30A outlet temperature.
Without being bound by any particular theory, some of the advantages of the system 702 include: the operating parameters in the first unit 30A can be modulated to employ higher amounts of TIOR materials and lower hydrogen purity heavy aromatics in order to enhance the liquid yield resulting from elevated direct incorporation of higher hydrogen-content feeds into the reactor feedstock; the second unit 30B can be operated with a vapor gap at the top, which can eliminate the need for a vacuum unit , facilitating the separation of ash from the reactor liquid as well as generating a high quality donor solvent containing an optimized Ash that can be integrated with one or more of the satellite processing units 604, 608.
As will be appreciated by those skilled in the art, the upgrading and processing that occurs in reaction unit 24 can use one or more difficult to process heavy-oils as a feedstock.
In some implementations of the present disclosure, the reaction unit 24 can receive a difficult to process heavy-oil feedstock from more than one source. For example, some implementations of the present disclosure can process a difficult to process heavy-oil feedstock from a primary upgrading facility, such as a coker-fractionator unit and one or more satellite upgrading facilities, such as a further a coker-fractionator unit, a visbreaker unit and/or a hydro-visbreaker unit. These implementations can improve the economics and decrease the greenhouse gas production of the primary upgrading facility and the satellite upgrading facilities.
Furthermore, one or more products of system 700 or system 702—such as the contents of either or both of conduit 370 and conduit 396, as discussed herein below—can be can be communicated with conduit 116A, which provides the product of the distillation system 101 to the heater 20. For example, the contents of conduit 370 can provide a high TIOR and/or Ash material to be coked, or gasoil boiling range polar aromatic donor stream containing excess Ash to be removed by coking within the coke drum 21. The Ash gets coked and the aromatic oil donor-solvent can reduce the coker unit 200 coke yield by hydrogen donation. The thermally processed donor solvent can also flash back into the coker-fractionator unit 200A to be recycled to either or both of the units 30A, 30B with low hydrogen-content coker fractionator bottom products or yielded as the 221 products and hydrotreated in a downstream coker hydrotreater.
In some implementations of the present disclosure, either or both of conduit 222 and conduit 224 can be routed to communicate with one or both of conduits 382, 386.
The contents of conduit 226 and conduit 228 can be excellent sources of polar aromatics and in some implementations of the present disclosure these conduits 226, 228 can be routed to communicate with the contents of conduit 300 or the contents thereof can be used as a carrier for additive make-up in the vessel 38.
In some implementations of the present disclosure, the additive can be directly added to liquid in the bottom of fractionator 22, thereby eliminating the need for vessel 38.
In some implementations of the present disclosure, part or all of the contents of conduits 360, 370 and 396 can be charged to fractionator 22 such that some or all products from the system 700 or the system 702 can be recovered at the coker fractionator with the coker products.
The fractionator unit 200A can also provide a source of a low-asphaltene solvent for co-processing difficult to process heavy-oil feedstock in either or both of system 700 or system 702. If additive is added to the coker fractionator, the additive will act to reduce coking in the fractionator bottoms.
Other Thermal Processors
Stream 422 and 424 can be routed to one or all of conduits 382, 386.
Stream 426 and 428 are excellent sources of polar aromatics and can be routed to 300 and/or used as a carrier for additive make-up in vessel 38.
In some implementations of the present disclosure, additive may be directly added to liquid in the bottom of fractionator 42, thereby eliminating the need for vessel 38.
In some implementations of the present disclosure, part or all of the contents of conduits 370 and 396 can be charged to fractionator 42 such that the ITP products can be recovered at the visbreaker fractionator with the visbreaker products.
Without being bound by any particular theory, the visbreaker unit 400A may provide the benefits of: directing the visbreaker fractionator tower bottoms for processing by one or both of the systems 700, 702 instead of processing into a low value product; protecting downstream fixed-bed catalysts; increasing visbreaker yields by providing substantially higher conversion, potentially in excess of 70% 975+ F conversion due to the supply of hydrogen—and anticoking ash from loops 600 or 612.
In some implementations of the present disclosure, the hydro-visbreaker unit 500A can include heater 53 for heating gas provided by conduit 534, for conducting at least a portion of the contents of the conduit 518, and the recycle gas within hydrovisbreaker soaker drum and separator system 51.
Conduit 522 and conduit 524 can be routed to communicate with the contents of one or both of conduits 382, 386.
The contents of conduit 526 and conduit 528 can be excellent sources of polar aromatics and can be routed to communicate with the contents of conduit 300 and/or used as a carrier for additive make-up in vessel 38.
In some implementations of the present disclosure, additive may be directly added to liquid in the bottom of fractionator 52, thereby eliminating the need for vessel 38.
In some implementations of the present disclosure, part or all of the contents of conduits 360, 370 and 396 can be charged to fractionator 52 such that the ITP products can be recovered at the visbreaker fractionator with the visbreaker products.
In some implementations of the present disclosure, the hydro-visbreaker unit 500A can also include a conduit 535 for conducting a purge of TIOR materials and/or Ash from the hydro-visbreaker soaking drum 51. This stream may be routed to conduit 300 and/or to be coked in a coker.
Without being bound by any particular theory, the hydro-visbreaker unit 500A may provide the benefits of: directing the hydro-visbreaker fractionator tower bottoms for processing by one or both of the systems 700, 702 instead of processing into a low value product; protection of downstream fixed-bed catalysts; an increased hydro Visbreaker yield improvement due to a substantially higher conversion potentially in range of 80% 975+ F conversion that is caused by the loops 600, 602 and the use of the anti-coking additives in either of the systems 700, 702; recycling the roughly 5-6 wt % gas yield generated by thermal conversion within conduit 518 through gas contacting loop with the heater 53 that can increase the temperatures of the gas to greater than 1000° F.; and, return of TIOR materials and/or Ash to one or both of systems 700, 702 for TIOR conversion, additive recycle and regeneration and aromatic oil donor-solvent recycle and regeneration. In some implementations of the present disclosure, the hydro-visbreaker unit 500A can provide a source of high-density, highly complex aromatic rings compound that efficiently convert from gas to liquid in either of the systems 700, 702. With the capabilities of systems 700, 702 to convert the contents of conduits 530 and 535, the hydrovisbreaker reaction system 51 can be designed with features of the reactor unit 30 and operated at pressures of less than 1000 psig.
In some implementations of the present disclosure, each of the upgrading systems 10, 10A and 10B can produce various gases and naphtha streams that can be conducted to either or both of the systems 700, 702 where the naphtha can be used as a diluent that can then be flash separated from one or more valuable products.
Upgrading Process
Some implementations of the present disclosure relate to a process 800 for upgrading difficult to process heavy-oil feedstocks that is performed by either of the systems 700, 702 described above.
Some implementations of the present disclosure relate to a process 800A that includes subjecting the difficult to process heavy-oil feedstock to a first step 812 of an ITP cracking process within an ITP reactor unit and then to a second step 814 of the ITP cracking process within a second ITP reactor unit. Each of the first step 812 and the second step 814 of the ITP cracking process will utilize a partial pressure of hydrogen that can be the same or that can be different between the steps. For example, the first step 812 of the ITP cracking can have a lower partial pressure of hydrogen than the second step 814.
As shown in
The second step 814 includes a step of conducting the liquid stream from the first step 812 into the second ITP reactor unit for a step 838 of generating further ITP cracking products within the second ITP reactor unit. As will be appreciated by one skilled in the art, many of the steps described for the first step 812 can also occur during the second step 814, including but not limited to the direct incorporation of light hydrocarbon feedstocks with high hydrogen content into the liquid stream that enters into the second ITP reactor for producing the further ITP products that have a volumetric gain, as compared to if there was no direct incorporation step. The further ITP products are then subjected to a step 840 of conducting towards a separation step 842 for separating a mixed effluent product from the second ITP reactor unit into a vapor stream and a liquid stream. The liquid stream can be subjected to a step 846 of conducting to a vacuum process and/or to the first reactor unit and/or the second reactor unit. Optionally, the hydrocarbon-rich vapor stream can be communicated with the mixed effluent product from the second reactor unit and/or the vapor stream from the separation step 842. The vapor stream from the separation step 822 is subjected to a step of removing impurities such as nitrogen and/or sulfur and the purified vapor stream can be subjected to a step of further separating into a hydrogen-rich vapor stream (which can be conducted upstream of the first reactor unit or not) and a hydrocarbon-rich vapor stream that includes the ITP products. This hydrocarbon-rich vapor stream can be subjected to a step for separating the ITP products into the various valuable constituent products.
The second step 814 can also include a step of introducing a gas stream with a higher partial pressure of hydrogen into the vapor stream from the separation step 822 prior to the step of removing impurities.
In some implementations of the present disclosure, an example of the volumetric gain achieved by using the fractionator tower bottoms as the heavy oil feedstock and using rich fuel gases as the low molecular-weight hydrogen feedstock with a medium and/or high hydrogen-content within an ITP reactor unit can be characterized as follows: when about 24 thousand barrels per day (kbpd) of heavy oil feedstock and about 21 kbpd of bitumen are conducted into the ITP reactor unit, through direct incorporation of the rich fuel gases (or other moderate and/or high hydrogen-content materials), there can be a total volumetric output of ITP cracking products of about 60 kbpd. This is about a 33% volumetric increase due to the direct incorporation of the rich fuel gases (or other C1 to C5 alkanes).
Data sets obtained during the upgrading of four different heavy oil feedstocks were analyzed and modeled for generating the data presented in the figures and for supporting the implementations of the present disclosure. For example, the data supports direct incorporation of low molecular-weight hydrocarbons into thermally processed, difficult to process heavy oils by a mass transfer of carbon and hydrogen atoms into valuable liquid products. One feedstock data set was obtained from a virgin, high sulfur asphaltic vacuum tower bottoms feedstock that was subjected to a pilot plant, slurry-phase hydrocracking unit that used a coal based anti-coking additive (pilot plant). A second feedstock data set was obtained from a virgin, high sulfur asphaltic vacuum tower bottoms feedstock (Asphalt VTB) that was subjected to a commercial slurry-phase hydrocracking unit that used an iron sulfide (FeS) additive system but did not include a gas-contacting system that provided the high mixing environment (the base commercial-operation or unit base). A third feedstock data set was obtained from (Asphalt VTB) but that was subjected to a commercial slurry-phase hydrocracking unit that used a FeS additive system and the gas-contacting system provided a high-mixing environment with a 73% partial pressure of hydrogen gas in the recycle including C1 through C6 low molecular weight components in the unit make-up gas (the modified commercial-operation or modified unit). A fourth feedstock data set was obtained from visbreaker bottoms of the Asphalt VTB (Visbreaker Bottoms Feed) that was subjected to the modified commercial-operation.
One alternative to the carbon rejection processes described above and the associated loss of liquid yield is hydrogen addition.
The hydrogen content of the reactor liquids shows the same increase in hydrogen content with decreased hydrogen partial-pressure and the use of the anti-coking additive A, as observed in the reactor deposits.
Without being bound by any particular theory, this pilot-plant data can indicate that higher hydrogen partial-pressure stabilized the cracked material within the light gases. At higher hydrogen partial-pressures, the hydrogen content of both the liquid products and reactor solids was lower than that achieved at lower hydrogen partial-pressures. The potential convertibility of the unconverted feed was decreased at higher hydrogen partial-pressures. The higher hydrogen partial-pressure environment resulted in a lower liquid yield of about 7.7 wt % of feed for the anti-coking additive A testing. The use of the anti-coking additive A improved the convertibility of the unconverted feedstock in all operations.
Superimposing the addition of hydrogen from water in the reaction mix resulted in intermediate levels of hydrogen in the liquid and reactor deposit products relative to the two hydrogen partial-pressure references discussed above. The hydrogen uptake from the gas charged to the closed pilot-plant system was reduced by about 0.1 wt %, based on the total gas charge, as shown on
The syn-gas pilot-plant work showed that increasing hydrogen partial-pressure in a closed, long residence-time environment resulted in a substantial loss of liquid yield and more hydrogen consumption due to the stabilization of carbon and hydrogen within the light gases.
When the reaction temperature was increased by about 18° F. and the amount of anti-coking additive B was increased by a factor of four, the net amount of methane incorporated at 2000 psig increased from about 100 SCFB to about 400 SCFB. The breakeven pressure, at which there was no net methane generated, decreased by about 200 psi and there was an elevated net methane-yield below about 800 psi at the elevated temperatures.
Similar to the syn-gas cracking pilot work, the pilot work done with methane gas was not a feasible commercial operation due to the large amount solids generated in the thermal process and deposited in the pilot plant. This pilot work does, however, demonstrate that methane can be directly incorporated in the reaction product mix at elevated operating pressures and the rate increases as a function of the system's pressure.
The complex nature of the products generated in cracking reactions performed at thermal conditions used in the upgrading of heavy oil can be shown by model-compound cracking studies. The cracking data shown in
To satisfy the hydrogen demand for the equilibrium on cracked materials generated in this closed system, aromatics were formed. The equilibrium carbon number distribution occurs around the C9 carbon, as show in
FCCUs can be operated with intermediate product cuts recycled into the thermal process with the result that the net yield of that material is substantially reduced and the product-cut can be eliminated. Similarly, high hydrogen content materials such as C5-C6 components of a conventional crude can be co-processed with a typical FCCU feedstock with the result that most of the C5-C6 components are incorporated into the typical FCCU product distribution. These results are consistent with the results shown in
Unlike the catalyzed FCCU system, the processing of Athabasca bitumen required extending the contact time with an associated increase in the C1-C2 yields.
As shown in
The olefinicity profile for the delayed coker C1-C4 is shown in
It is known to co-process bitumen VTB type feedstocks with a 400-630° F. distillate boiling range hydrogen donor solvent prepared in a conventional fixed-bed catalyst system. The external hydrotreater generates the donor solvent by saturated naphthalene to tetralin. In the bitumen VTB thermal processing environment, the narrow boiling range distillate containing the tetralin converted back to naphthalene and donated hydrogen to the thermal reactor liquid. The donation of the hydrogen in the liquid phase reportedly extended the potential 975+ F liquid conversion up from the visbreaker range to about 70% without the generation of coke. A distillate boiling range hydrogen donor enables hydrogen addition to a heavy oil at substantially reduced upgrading operating pressures.
In some implementations of the present disclosure the ITP is different from a SHC process. Such differences at least include increasing the localized heat input and reaction temperature at the point of the inlet gas. Contrary to conventional wisdom, these differences can result in a substantial reduction of polar aromatic partitioning during the thermal upgrading process. Further, by reducing the hydrogen purity in the hydrogen rich gas recycle stream, there was a substantial improvement in the reduction of polar components contained in the 975+ F liquid in the SHC products.
In contrast, the data for the modified commercial-operation showed a moderate increase in the nitrogen and, therefore, the degree of partitioning of the asphaltenes is greatly reduced.
The FeS based anti-coking additive system was common to both the base commercial-operation and the modified commercial-operation (shown as the triangle data-points in
The hydrogen content of the 975+ F product from the modified commercial-operation did not decrease with an increasing 975+ F conversion of the feed. At the 90% 975+ F conversion reference, the hydrogen content of the 975+ F liquid was about 2.8 wt % higher than the base commercial-operation. The hydrogen content of the unconverted 975+ F liquid increased as the molecular weight of the 975+ F product decreased associated with an increased 975+ F conversion in the modified commercial-operation.
Without being bound by any particular theory, the reduced polar aromatic partitioning observed in
The pilot plant operation was used to process the Cold Lake vacuum tower bottoms (VTB) using a coal-FeS anti-coking additive. The pilot plant feedstock is represented by the 1.050 specific gravity Cold Lake Analysis (see
The base commercial-unit was also used to process the Cold lake VTB using the FeS anti-coking additive. The base commercial-operation feedstock is represented by the 1.041 specific Cold Lake analysis (see
The modified commercial-unit was used to process VTB from an asphaltic crude. The asphaltic crude VTB that was processed over the period of the data set had a feed sulphur of about 4.0 wt %. Based on the source of the crude, the nC7 asphaltene content was in the range of about 28 wt % with a resin to asphaltene ratio of about 0.7. The hydrogen content of asphaltic VTB would be similar to the 10.5 wt % in the 1.041 specific gravity Cold Lake operation benchmark.
The modified commercial unit was also used to process visbreaker bottoms that were derived from asphaltic crude VTB and, therefore, the visbreaker bottoms were lower in hydrogen content than the asphaltic VTB charged to the visbreaker. The three samples of visbreaker bottoms shown in
The impact of the hydrogen partial-pressure is shown in
As shown in
The methane yields shown in
The C2 and C3 yields shown in Cold Lake pilot-plant work followed the same relative shifts as observed in
The data of
The reduction in the yield of the C1-C4 hydrocarbons achieved in the modified commercial-operation resulted in mass being transferred into the C5+ liquid yield.
The reduction in the C1-C4 yields and the increase in the hydrogen content of the 975+ F liquid is primarily due to the saturation and alkylation of the higher hydrogen content molecules into these highly condensed aromatic rings. The greater the number of fused aromatic rings, the greater the rate of saturation and alkylation similar to the relationship shown in
As the asphaltenes undergo thermal decomposition, low hydrogen content Toluene Insoluble Organic Residues (TIOR) components are generated. If allowed to self-associate, TIOR will generate coke.
The pilot plant data shown in
In contrast,
There are a number of process variables that are available to manipulate the TIOR:Ash ratio. As the TIOR:Ash ratio increased, agglomeration causes the particle size to increase. The FeS-based ash is about 5 times the density of the liquid within the reactor and the agglomeration of the ash with the TIOR can result in large particles that are readily gravity separated by differential settling velocities. The resulting differential settling velocities provided a mechanism for the segregation of the TIOR from the bulk reactor solution and for transporting the TIOR to the gas contactor at the bottom of the reactor. This behaviour is analogous to a clarification process for segregating asphaltenes to the bottom of a clarifier. This can provide a mechanism to concentrate and position the highly complexed aromatic ring clusters in the proximity of the high temperature gas contactor jets. This transportation of the TIOR through complexing with the dense Ash can facilitate a rapid alkylation of the high hydrogen-content gases into the aromatic ring clusters.
One of the characteristics of a contacting device is that a gas-contacting zone is created around the gas contactor where the TIOR-Ash complex is concentrated. A reactor withdrawal point at the center of the reactor about 3 feet above the gas contactor has been found to be a highly effective method for withdrawing Ash and any associated TIOR from the reactor. Additionally, this concentration of the TIOR and Ash creates an effective environment for maximizing the direct incorporation of the high hydrogen-content hydrocarbons into the concentrated highly aromatic oil. The ability to control the Ash accumulation in the reactor and the solubility of the TIOR by manipulating the polar aromatic support environment provides a mechanism to control the TIOR-Ash complex settling velocity. Precipitation and hydrogen transfer efficiency is controlled through the injection of polar aromatic solvents below the gas-contacting system. However, the addition of higher hydrogen content high boiling feedstocks such as paraffinic VTB in the top half of the reactor can promote flocculation and concentration of the TIOR and Ash into the gas-contacting zone at the bottom of the reactor. This provides both more hydrogen content in the overall feed to the unit and more multi aromatic clusters that can incorporate light gases.
As outlined in the preceding sections, there could be necessary trade-offs in terms of the need to control TIOR generation and prevent coke generation during the thermal upgrading of asphaltenes being balanced against the loss of liquid yield associated with the stabilization of low carbon number saturates at elevated hydrogen partial-pressures. Increasing the hydrogen transfer rate with increased hydrogen partial-pressure results in the production of undesirable low carbon number saturate products at the expense of a substantial reduction of the more valuable C5+ liquid products. In the thermal processes, increasing the contact time or increasing the bulk solution temperature further reduces the production of the more valuable C5+ liquid yield.
Implementations of the present disclosure relate to a mass transfer of higher molecular-weight light gases to overcome the polar partitioning associated with the thermal upgrading of asphaltenes. The hydrogen content of the TIOR and the asphaltenes are increased by reactions that directly incorporate higher hydrogen-content gases into the thermally generated heavy oil intermediate product. This process can allow the feedstock to be upgraded without generating undesirable very low hydrogen content residual fuel-oil. The higher hydrogen-content gases, such as the rich fuel gases, are injected at elevated temperature, for example in excess of about 800° F., 900° F. or 1000° F. creating a very localized environment at the contactor-bulk solution interface. In some implementations of the present disclosure, the higher hydrogen content gases are injected at between about 850° F. and about 1100° F. or between about 890° F. and 1000° F. with a velocity of at least about 200 feet/section (ft/s) or between about 250 ft/s and about 550 ft/s or between about 300 ft/s and about 500 ft/s. In some implementations of the present disclosure, the molecular weight of any H2 and the light hydrocarbon feedstock is at least about 2 pounds per mole, at least about 5 pounds per mole, at least about 10 pounds per mole, at least about 15 pounds per mole or greater. At these conditions of elevated temperatures, velocities and molecular weights of the injected gas, the cracking reaction can favour olefin formation. These olefins readily alkylate with the multi-aromatic structures within the TIOR and to a lesser extent with the polar aromatic co-processed solvents that stabilize the complex aromatic structures. Elevated jet temperatures exiting the gas contactor and the direct injection of olefins in the gas stream entering the reactor promote the alkylation of the light hydrocarbons onto the aromatics and the stabilization of the aromatics with elevated hydrogen to carbon ratio. Similar to the FCCU model compound example for octadecane, a very large number of reactions occur very quickly as new equilibriums are established, and the bulk reactor-temperature rapidly quenches the energy input from the gas contactor and the hydrocracking heat that is released.
In some implementations of the present disclosure, reducing the polar partitioning can reduce the need for high hydrogen partial-pressures to control the TIOR yield. As the asphaltenes and TIOR are being upgraded, the radicals are capped by alkylation with the high hydrogen-content gases rather than hydrogen because of the lower hydrogen concentration. The process can retain the thermally created olefins longer, which can enable these olefins to participate in the process and thereby generate less methane, ethane and propane yield as the vapours pass upward through the reactor.
During one specific test, the reactor temperature was increased to 873° F. with the average temperature over the day maintained at 868° F. The increase in operating temperature resulted in the conversion of all the 975+ F feed such that the liquid and ash stopped overflowing the reactor and a vapour space was generated and maintained at the top of the reactor. The gas make-up to the reactor consisted of about 20% C1+ hydrocarbons with the remainder being hydrogen typical of a fuels naphtha reforming process. The maximum temperature that the reactor could be run at was limited by the ability to maintain pressure on the reactor as the gas supply volume was limited and all the make-up gas was being incorporated in the liquid product. The elimination of the purge gas occurred even with elevated thermal conditions in the reactor. The average TIOR:Ash ratio for this operation is shown at 1.0 on
The modified commercial-process experienced a noticeable increase in exothermic reactions as the average reactor temperature was increased. Generally the hydrocarbon feed preheater outlet temperature was increased to increase the reactor temperature. However at about 840° F., reducing the hydrocarbon furnace outlet temperature was necessary to start. As the reactor temperature is increased, the energy input into the hydrocarbon furnace has to be set to at minimum value or eliminated. The controlled exothermic behaviour of the reaction is another advantage in reducing the overall CO2 generation with the process.
Table 1 below shows the hydrogen balance between the lower boiling point cut and the higher boiling point cut.
The lower hydrogen content in the 1256+ F cut yielded a higher C1 to C4 yield with the hydrogen contained in the C1-C4 yield being about 1.99 wt % and about 2.31 wt % for the 1256− F and the 1256+ F cuts, respectively. The C5+ liquid yield for both feedstocks was essentially the same at 10.7 wt % of the total product.
This example shows that the base commercial-operation is capable of upgrading nC7 asphaltenes. The inclusion of the nC7 asphaltenes in the SHC feed resulted in a higher density and lower hydrogen content feedstock. When upgraded, the higher density feedstock yielded about 14.5 vol % more liquid yield than the lower boiling VTB cut where the nC7 asphaltenes had been removed. Processing these feedstocks with implementations of the present disclosure could be beneficial to the low hydrogen-content feedstock. Not only is there a greater C1-C4 yield to be reduced, there are more large aromatic molecules and asphaltenes in the feedstock that can provide a base for alkylating any higher hydrogen-content materials that can be introduced through the gas contactor. Effectively, the direct incorporation of the high hydrogen-content hydrocarbons can result in essentially the same decreased amount of chemical hydrogen input (for example from a reformer) to achieve similar feedstock conversion and product qualities. However, the C5+ liquid yield would be substantially higher for a higher density feedstock.
Number | Date | Country | Kind |
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3011027 | Jul 2018 | CA | national |
Number | Date | Country | |
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Parent | 16509337 | Jul 2019 | US |
Child | 17836849 | US |