Patent Grant
7618530
The present invention relates to a method for the production and use of hydrogen donor solvents to increase heavy oil-to-hydrocarbon distillate conversion efficiency.
Terminology is important, especially for a complex field like hydrocarbon processing that progressed in parallel and very non-linear scientific and engineering practice pathways. Originally, heavy oils were hydrocarbons with a high density for a given boiling point range. However, the term heavy oil is often used interchangeably with high boiling by practicing engineers because most oil fractions with higher densities also have higher boiling points. However, some highly paraffinic oils or oil fractions may have significantly higher boiling points than much heavier, i.e., denser, aromatic oils or oil fractions. For the purposes of this invention, a heavy oil contains a significant quantity of a high density vacuum residual oil. Residual oils, also called residua or resids, are typically those fractions which are non-distillable under given conditions and remain at the bottom of a vacuum distillation tower and have equivalent normal boiling point (NBP) greater than approximately 525° C.
The efficiency of processes to convert heavy oils to distillates is generally determined by the relative rates of cracking reactions to produce lower molecular weight species and the rate of free radical polymerization reactions to produce higher molecular weight and less soluble species. The polymerization reaction rate dramatically accelerates to form solid petroleum coke when the polymerization reaction products form a separate mesophase. Therefore, control of both the heavy oil conversion and process solvent and solute properties are important.
The solubility of residual oil components in alkanes (paraffins), e.g., propane, butane, pentane, hexane, and heptane) has been used by petroleum refiners to up-grade residual oils and by researchers to obtain more detailed information about these component properties. A two-product commercial deasphalting unit produces deasphalted oil (DAO) and asphaltene streams and a three-product commercial deasphalting unit produces DAO, resin, and asphalt streams. The DAO, resin, and asphalt stream properties vary over a wide range depending on the deasphalter operating conditions. Broadly, the deasphalter product aromaticity and molecular weight have the following ranking: DAO<resin<asphaltene. Petroleum chemists use similar terms, with substantially different meanings, to specify residual oil solubility classes. For petroleum chemists, asphaltenes and maltenes are terms used to describe the insoluble and soluble fractions of a vacuum residue or deasphalter asphalt product. They are defined by the respective insolubility and solubility of these fractions in light hydrocarbons such as n-pentane, n-hexane, or n-heptane. As a result, pentane-insoluble-asphaltenes would have a lower molecular weight and aromaticity than heptane-insoluble-asphaltenes. The petroleum chemists usually define oils and resins as maltene species that readily adsorb on a packing and can be readily desorbed using alkane and polar solvents, respectively.
This invention defines (1) coke precursors as marginally soluble species in the heavy oil conversion process solvent and (2) process solvent properties in terms of the equivalent deasphalter DAO and asphaltene residual product properties. Theoretical methods to estimate solubility can be used to analyze heavy oil conversion process data [Jianzhong Wu, John M. Prausnitz, and Abbas Firoozabadi, “Molecular-Thermodynamic Framework for Asphaltene-Oil Equilibria”, AlChE Journal, Vol. 44, No. 5, May 1998].
This invention provides a heavy oil-to-distillates conversion method that is differentiated from and superior to the related art. This task is complicated by the fact that both this invention and the related art utilize very complex and poorly understood thermal cracking reactions to convert heavy oil to distillates. Molecular weight, elemental analysis, nuclear magnetic resonance (NMR) spectroscopy, and X-ray diffraction (XRD) analyses can be used to estimate the average structural data for hydroconversion feed and products [George Michael, Mohammad Al-Siri, Zahida Hameed Khan, and Fatima A. Ali, “Differences in Average Chemical Structures of Asphaltene Fractions Separated from Feed and Product Oils of a Mild Thermal Processing Reaction,” Energy Fuels, Vol. 19, No. 4, pages 1598-1605, 2005]. Even these very time consuming and expensive analytical methods provide only very general guidance to assess process performance. As a result, heavy oil process developers are forced to use less rigorous and costly methods to characterize, evaluate, and improve heavy oil conversion processes. Process development teams tend to use somewhat different approaches to analyze their processes and assess their performance relative to alternative approaches. The present inventor has discovered that the very general reaction system on
More specifically, the conversion process is initiated by thermal cracking of a carbon-carbon bonds (R—R′, where R and R′ represent the feedstock structure on either side of the ruptured bond) via Reaction 1 to form short lived free radical intermediate species (R. and R′.). These unstable free radical species can react with labile hydrogen atoms in the heavy oil to produce the desired stable reaction product via Reaction 2 (or Reaction 6). The labile hydrogen is typically a naphthenic hydrogen atom that is bonded to a carbon atom that is in the alpha position relative to an aromatic carbon (see hydrogen donor diluent example on
Visbreaking is a well-known petroleum refining process in which heavy oils are thermally cracked, under comparatively mild conditions, to provide products having lower viscosities and pour points to reduce the amount of less-viscous and more valuable blending oils required to produce a fuel oil product. Early visbreaking processes typically heated heavy oil in a fired heater to between 825° F. (441° C.) to 900° F. (482° C.) at moderate pressure with the maximum resid oil converted to distillates is limited by coke formation. More severe cracking conditions are not possible because of excessive coke precursor and coke formation. U.S. Pat. No. 2,762,754 taught that the maximum resid conversion could be significantly increased by increasing the visbreaker operating temperature from 900° F. (482° C.) to 1000° F. (538° C.). The rate of thermal cracking reactions (Reaction 1) increase much more rapidly with increasing temperature (activation energy≈45 kcal/mole) than free radial polymerization reaction rate (activation energy≈5 kcal/mole). As a result, it is not surprising that the maximum resid conversion increases with increasing operating temperature.
U.S. Pat. No. 2,843,530 further increased the maximum resid conversion by thermal cracking of the heavy oil in the presence of a hydrogen donor diluent. The hydrogen donor diluent was produced by catalytic hydrogenation of an aromatic distillate stream comprising thermal tars, catalytic cycle stocks, and lube oil extract. The subsequent hydrogen donor diluent cracking (HDDC) process development effort focused on HDDC process improvements and more cost effective hydrogen donor solvent production and regeneration methods.
The hydrogen donor diluent cracking (HDDC) process development effort focused on the use of additives, optimization of HDDC process operating conditions and pre-treatment of the feeds. U.S. Pat. No. 2,873,245 teaches a two-stage HDDC process. U.S. Pat. Nos. 2,989,461; 4,389,303; and 4,592,830 teach the addition of molecular hydrogen to the HDDC process feed to increase the maximum resid conversion. U.S. Pat. No. 4,587,007 teaches the addition of thiols to the HDDC process feed. U.S. Pat. Nos. 4,454,024; 4,487,687; and 4,485,004 teach the addition of molecular hydrogen and fluid catalytic cracking catalyst, coke solids, and hydrotreating catalyst, respectively, to the hydrogen donor diluent cracking process feed. U.S. Pat. No. 4,698,147 teaches a method to further increase the maximum operable resid conversion by operating at high temperature [>900° F. (482° C.)], low pressure [<1100 psig (75.8 bar)], and sufficient residence time to achieve the desired resid conversion.
U.S. Pat. No. 4,002,556 further increased the hydrogen donor diluent process efficiency by introducing the hydrogen donor diluent at multiple locations to optimize the rate of hydrogen transfer. U.S. Pat. No. 4,363,716 teaches recycle of a portion of the unconverted resid to increase the overall resid conversion. U.S. Pat. Nos. 4,451,354 and 4,514,282 teach hydrotreating the resid feed and recycle resid in the presence of a supported catalyst prior to treatment in a hydrogen donor diluent cracking process. U.S. Pat. Nos. 6,183,627 and 6,274,003 teach deasphalting the fresh feed and recycle resid feeds to the HDDC process. U.S. Pat. Nos. 4,347,120 and 4,604,186 teach methods to further increase the overall resid conversion by feeding the unconverted resid from the HDDC process to a delayed coker. U.S. Pat. No. 4,115,246 used partial oxidation of the unconverted resid from the HDDC process to produce a synthesis gas for hydrogen production. U.S. Pat. Nos. 3,238,118 and 4,363,716 teach a method to use a distillate hydrocracking unit to produce the distillate hydrogen donor diluent. U.S. Patent Application US 2003/0129109 teaches a method to produce the hydrogen donor precursor via thermal cracking. U.S. Pat. No. 4,090,947 teaches a method to use a premium coker gas oil as the hydrogen donor precursor.
Although, remarkably efficient heavy oil hydrogen donor diluent cracking processes have been developed over time, no commercially attractive approach has been identified to produce or regenerate the hydrogen donor diluent feed. As a result, there was a strong commercial incentive to develop a single-step heavy hydrocracking process.
U.S. Pat. No. 2,987,465 first introduced the ebullated bed hydrocracking reactor concept. The expanded catalyst bed was much less susceptible to plugging problems associated with heavy oil hydrocracking than the previous fixed catalyst bed designs. However, this design also had a major disadvantage: only the feed oil was available to expand the catalyst bed, which required using either inconveniently large reactor height to diameter ratio or small and difficult to separate catalyst particles. U.S. Pat. No. 3,207,688 eliminated this problem by adding a gas-catalyst-oil disengagement zone and oil recycle line. Virtually all the modern heavy oil hydrocrackers are based on this general design concept with many other mechanical and process improvements.
However, the heavy oil hydrocracker concept also has a significant problem: The resid thermal cracking reactions (Reaction 1 on
The rate of the thermal cracking reactions (Reaction 1 on
Fortunately, the hydrogenation reaction rate to the thermal cracking reaction rate ratio required to maintain reactor operability also decreases with increasing temperature (U.S. Pat. No. 4,002,556). Unfortunately, the actual hydrogenation reaction rate to thermal cracking reaction rate ratio decreases more rapidly than the required ratio to maintain reactor operability (U.S. Pat. No. 4,427,535). Therefore, heavy oil hydrocrackers have a maximum operating temperature that is a function of the catalyst hydrogenation activity and feedstock properties, primarily the concentration and effectiveness of hydrogen donor species.
U.S. Pat. No. 4,427,535 first faced this problem by teaching that the ebullated bed hydrocracker operating temperature must be limited such that the percent Ramsbottom carbon residue conversion is greater than the percent resid conversion to distillates to ensure successful operation of an ebullated resid hydrocracker. As a result, heavy oil hydrocracker development efforts have focused on methods to either remove coke precursors or increase the rate of hydrogenation to increase the maximum resid conversion and process efficiency.
U.S. Pat. No. 4,495,060 teaches the use of a rapid hydrocarbon quench of the ebullated bed hydrocracker liquid product to minimize coke formation in the product recovery system. U.S. Pat. No. 4,411,768 teaches removal of coke precursors from recycle resid feed to an ebullated bed resid hydrocracker by cooling the recycle resid, allowing the coke precursor to form a separate phase, and separating the coke precursor phase. U.S. Pat. No. 4,457,830 teaches the use of acids to remove coke precursors from recycle resid feed to an ebullated bed resid hydrocracker. U.S. Pat. No. 4,686,028 teaches the use of solvent extraction to selectively removal deasphalted oil from the resid hydrocracker to increase asphaltene solubility in the resid hydrocracker and convert the DAO to distillates more efficiently in either fixed bed hydrocracking or fluid catalytic cracking processes.
The related art has identified a wide variety of methods to increase the hydrogenation rate and the heavy oil hydrocracker maximum operable temperature and resid conversion. U.S. Pat. Nos. 4,640,765; 4,686,028; and 5,980,730 teach that the addition of a hydrogen donor solvent, deasphalter resin fraction and deasphalter DAO, respectively, to the ebullated bed hydrocracker feed increase reactor operability. U.S. Pat. No. 5,932,090 teaches the use of fine catalyst to increase the rate of hydrogenation in an entrained flow reactor with catalyst recovery and recycle. U.S. Pat. No. 5,362,382 teaches a two-stage heavy oil process, in which the first stage operates at milder conditions than the second stage. U.S. Pat. Nos. 5,164,075 and 5,288,681 teach methods to produce colloidal heavy oil catalysts that are particularly effective for hydrogenating asphaltenes. WO 2004/056946, WO 2004/056947, and U.S. Pat. Nos. 5,294,329; 5,298,152; and 6,511,937 teach methods to recover and recycle colloidal heavy oil hydrocracking catalysts. U.S. Patent Application US 2005/0241993 teaches the addition of colloidal hydrotreating catalyst to an ebullated heavy oil hydrocracker and operating the reactor gas-liquid separator within 20° F. (11° C.) of the hydrocracker temperature to decrease the rate of coke precursor formation.
Clearly, both the heavy oil hydrogen donor diluent cracking (HDDC) and hydrocracker processes have been subject to intensive and innovative development programs. However, this extensive effort has failed to find a commercially attractive approach to produce the hydrogen donor diluent or thermal crack the heavy oil under optimum conditions.
The present invention provides for converting heavy oils by using a resid hydrocracker or resid hydrotreater reactor to produce hydrogen donor solvent feed for a hydrogen donor cracking process with both steps operating at optimum operating conditions.
More particularly, the present invention provides for a method for hydroconversion of a heavy oil comprising
The present invention further defines a method for hydroconversion of a heavy oil comprising:
The first reaction zone utilizes conventional particulate and/or colloidal resid hydrogenation or hydrocracking catalysts. A conventional nickel-molybdate or cobalt-molybdate on alumina catalyst with a large pore size distribution is used as a resid hydrocracking or hydrogenation catalyst to maximize access of the large resid molecules to the catalyst surface. A conventional molybdenum disulfide colloidal catalyst may be advantageously used to facilitate hydrogenation of the resid in the first reaction zone and facilitate hydrogen transfer in the second reaction zone in step d. The temperatures and pressures at which the steps of these inventions are run are typically about 370° to 470° C. for step b at a hydrogen partial pressure of about 1000 to 3000 psig. At lower temperatures in this range (about 370 to 425° C.), a fixed bed, down-flow resid hydrotreater reactor may be advantageously used. An ebullated bed resid hydrocracker may advantageously be used throughout the step b temperature range (about 370° to 470° C.). In addition, the ebullated bed resid hydrocracker can advantageously use nickel-molybdate or cobalt-molybdate on alumina catalysts with a smaller particle size than the fixed bed, down-flow resid hydrotreater reactor. For most resid feedstocks, the higher temperature operation with an ebullated bed resid hydrocracker is preferred. As a result, the detailed process description will focus on the ebullated bed resid hydrocracker case and note adjustments required for the fixed-bed, down-flow resid hydrotreater option. The residence time in step b ranges from about 5 to 60 minutes.
In step d, the residence time would typically range between 0.01 and 100 seconds. The pressure during step d is between about 5 and 1000 psig. The use of a residual oil hydrogen donor solvent, rather than the conventional distillate hydrogen donor diluent, decreases the step d minimum pressure and hydrogen donor cracking reactor volume and eliminates the requirement to recycle a distillate hydrogen donor diluent precursor. The colloidal catalyst that is added to step a is entrained with the liquid product from step c and facilitates hydrogen transfer in the step d hydrogen donor cracking process.
The present invention further provides for a method for step d comprising
The process combining heavy oil hydrocracking or hydrotreating and hydrogen donor conversion process will be described with the aid of
The ebullated bed hydrocracker reactor 5 typically operates with a hydrogen partial pressure between 1000 and 3000 psig and a temperature between 370 and 470° C. As noted earlier, a fixed-bed, down-flow resid hydrotreater reactor may be employed at the lower range of these temperatures (about 370 to 425° C.). The heavy oil residence time in the ebullated bed hydrocracker reactor 5 is adjusted such that the quantity of hydrogen added to the oil meets or exceeds the requirements of the subsequent hydrogen donor cracking process step 12. The residence time is typically about 5 to 60 minutes. The residence time for both the ebullated and fixed bed reactor is conveniently estimated using the ratio of the catalyst bed volume to the heavy oil volumetric feed rate. The hydrogen donor cracking process step 12 typically has a hydrogen requirement equivalent to 100 to 500 standard cubic feet of hydrogen per barrel of resid hydrocracker feed heavy oil 1. The standard cubic foot measurement is determined at one atmosphere absolute pressure and a temperature of 60° F. Traditionally, a 42 gallon (≈159 liter) barrel is used in this determination. A recycle heavy gas oil 13 hydrogen donor precursor can be advantageously fed to the ebullated bed hydrocracker reactor 5 to facilitate the production of an appropriate hydrogen donor cracking process feed 14. The product distillation system 15 is operated to provide the maximum practical normal boiling point end point, typically between 500 and 535° C., for the recycle heavy gas oil 13 hydrogen donor precursor. The initial normal boiling point of the recycle heavy gas oil 13 hydrogen donor precursor is adjusted to provide the desired ratio of distillate-to-resid ratio in the hydrogen donor cracking process feed 14 stream.
The ebullated bed hydrocracker reactor 5 product 16 is separated into a vapor stream 17 and hydrogen donor cracking process feed 14 in a high pressure separator 18. The high pressure separator 18 is operated with a temperature that is essentially equivalent to the ebullated bed hydrocracker reactor 5 operating temperature and minimum liquid residence time to minimize fouling in the high pressure separator 18 and downstream equipment. Since the resid hydrocracker 5 typically operates at a substantially higher pressure than the hydrogen donor 22 cracker, the gravity vapor liquid separator 18 may be advantageously replaced by a cyclone separator to decrease the liquid residence time. A recycle hydrogen stream 19 and a light oil stream 20 are typically produced in the high pressure hydrogen recovery system 21 by condensation.
Typically, the hydrogen donor cracking process feed 14 has a 524° C.− distillate to 524° C.+ resid mass ratio between 0.1 and 2. The hydrogen donor cracker 12 comprises a heating furnace 22 and optional subsequent reactor volume 23, often called a soaking drum. The hydrogen donor cracking process pressure is typically between 100 and 1000 psig.
The conventional product distillation system uses moderate pressure and vacuum distillation to recover the gas 26, distillate 27, heavy gas oil (13 and 24) and heavy oil 28 products from the light oil 20 and hydrogen donor cracking process product 25. The heavy oil 28 product may contain spent colloidal catalyst that should either be recycled to the ebullated bed hydrocracker reactor 5 with the colloidal catalyst or the colloidal catalyst should be recovered from heavy oil 28 product and recycled via stream 2. For the purpose of this invention, the overall resid to distillate conversion is defined as unity minus the ratio of the mass of species with normal boiling points greater than 525° C. in stream 28 divided by the mass of the species with normal boiling points greater than 525° C. in stream 1.
In fact, it is very difficult to design a heating furnace 22 that can achieve the heating rates implied by
An annular Laval type convergent-divergent nozzle 46 is positioned down-stream of the annular combustion chamber 45. The combustion chamber 45 pressure is between 2 and 15 times the pressure in the hydrogen donor cracking reactor 23. The hydrogen donor cracking reactor 23 typically operates between 5 and 1000 psig. A hot and high velocity annular gas jet 47 is produced. The fuel 41 and oxidant 37 flow rates are adjusted to ensure an oxidant deficiency of between 2 and 10% in the annular gas jet 47. The hydrogen donor cracking process feed 14 is intimately mixed and rapidly heated by the annular gas jet 47. The oil is heated to between 500° C. and 800° C. for a residence time between 0.01 and 100 seconds to achieve the required resid conversion. The residence time in hydrogen donor cracking reactor is conveniently estimated as the ratio of the reactor volume to the heavy oil feed rate 14. The hydrogen donor diluent cracker 23 product 25 is readily cooled to less then 400° C. using a recycle heavy gas oil quench 27 to minimize formation of a separate asphaltene phase and form coke. The hydrogen donor diluent cracker 23 product 25 is purified using a conventional distillation system 15.
This invention is particularly useful for the production of heavy oils and bitumen.
The operation of the hydrogen donor cracking reactor 23 with the raw bitumen 48 and hydrogen donor diluent 58 blend is essentially equivalent to the operations with the ebullated bed hydrocracker heavy oil product 14. The major differences arise from integration of the hydrogen donor cracking reactor 23 and steam 51 production. A pump 63 is used to circulate the quench oil 24 through heat exchanger 64 to produce steam 51 for the bitumen production facility 50. A gas-liquid separator 65 removes the gaseous cracked products and combustion products 66 from the quench oil 24. A conventional steam boiler uses the balance of the synthesis gas 56 and low sulfur fuel oil 62 to produce the balance of the steam 51 requirement for the bitumen production facility 50.
While this invention has been described with respect to particular embodiments thereof, it is apparent that numerous other forms and modifications of the invention will be obvious to those skilled in the art. The appending claims in this invention generally should be construed to cover all such obvious forms and modifications which are within the true spirit and scope of the present invention.
| Number | Name | Date | Kind |
|---|---|---|---|
| 2620293 | Blue et al. | Dec 1952 | A |
| 2762754 | Offutt et al. | Sep 1956 | A |
| 2843530 | Langer, Jr. et al. | Jul 1958 | A |
| 2859169 | Herman | Nov 1958 | A |
| 2873245 | Thompson et al. | Feb 1959 | A |
| 2987465 | Johanson | Jun 1961 | A |
| 2989461 | Eastman et al. | Jun 1961 | A |
| 3207688 | Van Driesen | Sep 1965 | A |
| 3238118 | Arey, Jr. et al. | Mar 1966 | A |
| 3487121 | Hallee | Dec 1969 | A |
| 4002556 | Satchell, Jr. | Jan 1977 | A |
| 4090947 | Satchell, Jr. | May 1978 | A |
| 4115246 | Sweany | Sep 1978 | A |
| 4347120 | Anderson et al. | Aug 1982 | A |
| 4363716 | Greene et al. | Dec 1982 | A |
| 4389303 | Simo et al. | Jun 1983 | A |
| 4411768 | Unger et al. | Oct 1983 | A |
| 4427535 | Nongbri et al. | Jan 1984 | A |
| 4451354 | Stuntz | May 1984 | A |
| 4454024 | Singhal et al. | Jun 1984 | A |
| 4457830 | Kydd | Jul 1984 | A |
| 4475592 | Pachovsky | Oct 1984 | A |
| 4485004 | Fisher et al. | Nov 1984 | A |
| 4487687 | Simo et al. | Dec 1984 | A |
| 4495060 | Abrams | Jan 1985 | A |
| 4514282 | Poynor et al. | Apr 1985 | A |
| 4587007 | Rudnick | May 1986 | A |
| 4592830 | Howell et al. | Jun 1986 | A |
| 4604186 | Lutz et al. | Aug 1986 | A |
| 4640765 | Kubo | Feb 1987 | A |
| 4686028 | Van Driesen et al. | Aug 1987 | A |
| 4698147 | McConaghy, Jr. | Oct 1987 | A |
| 5164075 | Lopez | Nov 1992 | A |
| 5201650 | Johnson | Apr 1993 | A |
| 5288681 | Gatsis | Feb 1994 | A |
| 5294329 | Kramer | Mar 1994 | A |
| 5298152 | Kramer | Mar 1994 | A |
| 5362382 | Heck et al. | Nov 1994 | A |
| 5932090 | Marchionna et al. | Aug 1999 | A |
| 5980730 | Morel et al. | Nov 1999 | A |
| 6183627 | Friday et al. | Feb 2001 | B1 |
| 6274003 | Friday et al. | Aug 2001 | B1 |
| 6511937 | Bearden, Jr. et al. | Jan 2003 | B1 |
| 20030129109 | Bronicki | Jul 2003 | A1 |
| 20050241993 | Lott et al. | Nov 2005 | A1 |
| Number | Date | Country |
|---|---|---|
| 866.979 | Nov 1978 | BE |
| 0 023 802 | Feb 1981 | EP |
| WO 2004056947 | Jul 2004 | WO |
| WO 2204056946 | Jul 2004 | WO |
| Number | Date | Country | |
|---|---|---|---|
| 20070158239 A1 | Jul 2007 | US |
| Number | Date | Country | |
|---|---|---|---|
| 60758330 | Jan 2006 | US |
