Supercritical hydroextraction of kerogen from oil shale ores

Abstract

Water is added to reduce the hydrogen consumption in kerogen conversion. The water is not stripped from oil shale ores, instead up to 10% w/w of water is added in the kerogen conversion at 400-475° C. The recycle solvent for conversion is changed to a high boiling point fraction of the oil, with a minimum boiling point of the order of 245° C. Such repeated recycling of high boiling fraction oil with a hydrogen-donor mid-distillate will remove additional sulphur, oxygen and nitrogen. It improves the quality of the final product oil. Most of the high boiling fraction of the oil will separate from the oil shale ores, at around 450° C. and 650 psig. Then it is reduced to 250° C. at 100 psig before distillation. Such column operates at around 10 to 250 psig, but preferably 50 psig. This improves thermal efficiency by avoiding the need to condense high boiling point vapors. The addition of water provides further enhancement in operating results. A consequence of recycling high boiling oil fractions is the repeated hydrogenation and hydrocracking of that fraction to increase the proportions of lower boiling fractions. This reduces the asphaltene content of the recycled and final product streams. Such is similar to hydro-visbreaking of bitumen, in which both the asphaltene fraction and the sulphur content can be halved, resulting in a product of 25-30° API gravity in a single pass. The recycling of high boiling fraction means approximately half of the product oil is repeatedly recycled. There is no recycle of the low boiling fractions. It is expected that the intractable asphaltene content of about 5% of the product oil will be in suspension under hydro-treating and hydro-cracking conditions. These are removed with the solids after a supercritical solvent wash.

Description

RELATED PATENT APPLICATIONS

This Patent application is a continuation-in-part of these previous U.S. patent application Ser. No. 09/490,254, filed Jan. 24, 2000, and titled APPARATUS AND METHOD FOR THE SUPERCRITICAL HYDRO EXTRACTION OF KEROGEN FROM OIL SHALE ORES; Ser. No. 10/247,868, filed Sep. 19, 2002, and titled SUPERCRITICAL HYDRO EXTRACTION OF KEROGEN AND AQUEOUS EXTRACTION OF ALUMINA AND SODA ASH WITH A RESIDUE FOR PORTLAND CEMENT PRODUCTION; and, Ser. No. 11/404,623, filed Apr. 13, 2006, and titled PROCESSING PLANT FOR PRODUCING STABLE PIPELINEABLE CRUDE OIL FROM KEROGENOUS OIL SHALES.

BACKGROUND OF THE INVENTION

1. Field of the Invention

The present invention relates to processes for extracting the kerogen bituminous matter from oil shale ores to produce a pipelineable crude oil, and more particularly to extraction processes and apparatus that depend on super-critical hydrogen-donating (H-donating) solvents applied to carbonaceous oil shale ores with high kerogen contents and low Fischer Assay yields.

The present invention further relates to processes for extraction of associated minerals (such as alumina and soda ash in the oil shale ores) by aqueous means, and a residue for production of Portland cement. More particularly, to processes and apparatus that depend on the use of supercritical hydrogen-donatives (H-donor) solvents to remove the oil, and to aqueous leach processes that remove the alumina and soda ash values at temperatures of 400° C. And in particular, to processes in which alumina and soda ash values are removed by precipitation/crystallization in order to maintain an equilibrium of such values from an aqueous leaching circuit in alumina production.

2. Description of the Prior Art

Known oil shale ores deposits are immense, they are found on all the major continents of the world in massive deposits. The presently estimated oil shale ores reserves exceed known crude-oil resources by orders of magnitude. The oil shale ores reserves in the United States alone are estimated to represent over seven trillion barrels of oil. Such reserves are concentrated in the Green River formation of Utah, Colorado, Wyoming, and also in the Devonian-Mississippian Eastern Shale Deposits between the Appalachian and Rocky Mountains.

Retort and solvent processes are conventionally applied to the job of extracting the kerogen from the oil shale ores. Retorting processes are divided into in-situ and surface types. All such conventional processes require large amounts of heat. Retorting especially requires expensively high temperatures up to 574° C., and the gaseous heat transfer media used in surface retorting also need very large processing vessels for efficient production.

Such heating of oil shale ores creates an environmental problem because some of the constituents are swelled and the whole is too large to put back into the pit from which it came. Such spent oil shale ores can have a volume that is 130% of the original shale ore. Therefore, easy disposal of the spent shale in the original mine is not possible.

Retort reaction times and conditions must be carefully controlled to avoid visbreaking or cracking the heavy oil into hydrocarbon products with molecular weights that are too low. If the retort reaction times and conditions get too far out of control, largely unusable residual carbon output increases. The important hydrocarbons bound in oil shale ores are collectively called kerogen and have a mixture of high molecular weight components.

Kerogen is conventionally converted to more convenient forms by heating it to 350° C., or higher, to yield a range of hydrocarbons with lower molecular weights, e.g., methane to light oil. Retorting processes normally operate near 500° C. Extended reaction time leads to conversion of primary bitumen products to other lower molecular weight products and residual carbon. Retorting also typically produces unacceptable environmental emissions, relatively low yields of bitumen and requires heavy water usage.

The challenges in the processing of oil shale ores include limiting the production of gas products like methane to enough to fuel the process, and keeping the production of unusable carbon residue at a minimum. Thus, it is without the conversion to secondary lighter products. Upgrading to the desired final products is more efficient in downstream processing.

Kerogen in oil shale ores is relatively insoluble in most organic solvents at or below their normal boiling points. But if the environmental pressure is increased to raise the boiling point to higher than 600° K, solvents like toluene will dissolve the kerogen. Solvent extraction separates oil shale ores from spent shale without vaporization. The converted hydrocarbon products result from dissolution under reaction conditions, e.g., heating the oil shale ores and a solvent to 380° C.-540° C. Sometimes hydrogenation is also needed for good conversion. In general, solvent processes have better yields than retorting processes.

Although there are a number of solvent process variations, none efficiently separate spent-shale particles from the solvent and bitumen. And apparently no prior art processes have thought to use supercritical pressures to keep the solvents in their liquid phases at temperatures that would otherwise cause them to boil away.

Pao, et al., describe in U.S. Pat. No. 4,737,267, the difficulties associated with trying to use supercritical toluene as an extractant. For example, such process did not address stability issues, potentially lower yields of “carbonaceous” oil shale ores, nor did it address the olefin content of the produced oil.

Oil removal from oil shale ores and its residue from retorting has been leached for recovery of alumina and other values as described in several patents including multi mineral products such as superior oil U.S. Pat. No. 3,821,353, issued Jun. 28, 1974.

In multi-mineral research for such products from oil shale ores, it has been a prerequisite that the oil shale ores be retorted generally at temperatures around 500° C. in order to effectively remove oil and also release the alumina and sodium values from subsequent aqueous leaching systems. Since alumina values release with soda ash from dawsonite (in oil shale ores) at temperatures around 370° C., good temperature control is needed to process oil shale ores from the Green River Basin of Colorado, Wyoming and Utah. Some processing temperatures exceed 600° C., which locks up the minerals especially when some coke residue is burnt to supply the heat of the retorting process which vaporizes the oil product.

However in the United States, and elsewhere, retorting processes are unlikely to be used due to environmental and political constraints. In general, the CO₂ production inherent in coking produces about 0.3T CO₂ per barrel of oil, uses up to three barrels of water, and has reclamation problems due to increased spent shale volume.

Known oil shale ores deposits are immense, they are found on all major continents of the world in massive deposits. The presently estimated oil shale ores reserves exceed known crude oil resources by orders of magnitude. The oil shale ores reserves in the United States alone are estimated to represent over seven trillion barrels of oil. Such reserves are concentrated in the Green River formation of Utah, Colorado, Wyoming, and the Devonian-Mississippian Eastern Shale deposits between the Appalachian and Rocky Mountains.

The Green River formation is known to contain about two million barrels of oil associated with several billion tons of alumina and soda ash. In fact, several million tons per year of soda ash is commercially produced which leaves behind kerogen and alumina/soda ash values in the oil shale ores dawsonite. U.S. Pat. No. 6,010,672, issued Jan. 4, 2000, and others, describe mining techniques to access soda ash values which may occur independent of the kerogen and dawsonite in the shale.

The supercritical solvent extraction described by Pao in U.S. Pat. No. 4,737,267, uses toluene to convert kerogen to solvent-soluble oil products at 380° C. to 540° C. almost completely at the supercritical pressure of over 500 psig, while more or less completely washing the carbon off the residue. Such makes the residue suitable for aqueous leaching out the alumina and soda ash values contained in the associated dawsonite of the Green River oil shale ores. Other ore bodies have alumina and molybdenum which are extractable by high temperature, high pressure, aqueous leach. The residue in kerogenous are bodies where calcium carbonate predominates (such as oil shale ores in Green River, USA; Julia Creek, Australia, and Estonia) is usable in the production of Portland cement.

However, U.S. Pat. No. 4,737,267, did not address the stability issues connected to the olefin content of the produced oil. Separate hydrotreating was essential for a stable pipelineable crude. In addition, lower yields of oil from oil shale ores resulted in particular from “carbonaceous” oil shale ores.

SUMMARY OF THE PRESENT INVENTION

Briefly, a process embodiment of the present invention is a method for producing pipelineable synthetic crude oil from oil shale ores including those including cannel coal. Water is added to reduce the hydrogen consumption in kerogen conversion. The water is not preliminarily stripped from oil shale ores as in other processes. Instead, up to 10% w/w of water is added in the kerogen conversion at 400-475° C. The recycle solvent for conversion is a high boiling point fraction of the oil, with a minimum boiling point of the order of 245° C.

Repeated recycling of high boiling fraction oil and a hydrogen-donor mid-distillate results in increases in the removal of sulphur, oxygen and nitrogen. Fewer contaminants means the quality of the final product oil is improved. Most of the Kerogen will be converted to oil by the high boiling fraction of the oil, and will be separated from the shale residue, at around 450° C. and 650 psig. Then the separated oil is reduced to 250° C. at 100 psig before distillation. Such column operates at around 10 to 250 psig, but preferably 50 psig. This improves thermal efficiency by avoiding the need to condense high boiling point vapors. The addition of water provides further enhancement in operating results.

A consequence of recycling high boiling oil fractions is the repeated hydrogenation and hydrocracking of that fraction to increase the proportions of lower boiling fractions. This reduces the asphaltene content of the recycled and final product streams. Such is similar to hydro-visbreaking of bitumen, in which both the asphaltene fraction and the sulphur content can be halved, resulting in a product of 25-30° API gravity in a single pass.

The recycling of high boiling fraction means approximately half of the product oil is repeatedly recycled. There is no recycle of the low boiling fractions. It is expected that the intractable asphaltene content of about 5% of the product oil will be in suspension under hydro-treating and hydro-cracking conditions. These are removed with the solids after a supercritical solvent wash.

So adding water reduces how much additional hydrogen is required to run the process. Both marine and lacustrine oil shale ores are such that some form of reaction with the carbonaceous residue on the shale clay fraction effectively reduces the need for supplementary hydrogen.

An advantage of the present invention is that a process for kerogen extraction from oil shale ores is provided that uses water to reduce the amount of hydrogen needed in other similar processes.

Another advantage of the present invention is that a process for kerogen extraction from oil shale with intrusions of cannel coal is provided that produces oil products far in excess of the ore's standard Fischer assays, e.g., 400 liters to the ton, versus a Fischer assay of 119 liters to the ton.

A further advantage of the present invention is that a process for kerogen extraction from oil shale ores is provided that produces higher yields of oil with reduced gas production.

Another advantage of the present invention is that a process for kerogen extraction from oil shale ores is provided that produces hydro-visbroken crude oil which is stable enough for conventional pipeline transfer to refineries.

An advantage of the present invention is that a process for kerogen extraction from oil shale ores is provided that allows for the recovery of aqueous solutions of alumina and soda ash by simple leaching.

A still further advantage of the present invention is that a process for kerogen extraction from oil shale ores is provided that produces a minimum of environmental contamination. In addition to calcium carbonate rich oil shale ores, the residue is suitable for the direct production of Portland cement.

Another advantage of the present invention is that a process for kerogen extraction from oil shale ores is provided that recycles its solvents and heat.

These and other objects and advantages of the present invention will no doubt become obvious to those of ordinary skill in the art after having read the following detailed description of the preferred embodiment as illustrated in the drawing figure.

IN THE DRAWINGS

FIG. 1 is a functional block diagram of an oil shale ores processing plant embodiment of the present invention that implements a process for kerogen extraction from oil shale ores;

FIG. 2 is a functional block diagram of oil shale ores processing plant embodiment of the present invention. Such implements a process for extraction of alumina, soda-ash, and other mineral values from the aqueous leach liquor after oil is removed from the oil shale ores. The residue is useful in the production of Portland cement;

FIG. 3 is a functional block diagram of another oil shale ores processing plant in a second process embodiment of the present invention that is modified from that shown in FIG. 1;

FIG. 4 is a functional block diagram of an oil shale ores processing plant in a third process embodiment of the present invention that uses more water and less hydrogen than those shown in FIGS. 1-3; and

FIG. 5 is a recycle diagram showing how input oil shale ores is added to a recirculating +245° C. carrier to produce the product oil output in a 1-ton test plant experiment of the process and system described in FIG. 4.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENT

An oil shale ores processing plant embodiment of the present invention is diagrammed in FIG. 1 and is referred to herein by the general reference numeral 100. The principle product produced is synthetic crude oil that is suitable for pipeline transportation. An oil shale ores input feed 102 is crushed by a size reducer 104. Alternatively, the input feed 102 can comprise cannel coal, aka, “candle coal”, a type of coal with large amounts of hydrogen.

A crushed oil shale ores flow 106 is mixed with a recirculating solvent and input to a water remover 108. Any waste water 110 is removed from the system. A recovered solvent vapor flow 112 is added, condensed for heat recovery and a slurry flow 114 is output. A heat exchanger (H/E) 116 provides a recovered-heat flow 118 and outputs an oil product flow 120. Hot oil product flow 122 is received from further down the process and is stripped of its elevated heat for recovered-heat flow 118, and the remainder is the cooler oil product flow 120. A kerogen converter 124 outputs a slurry 126. Reaction gases are drawn off in a flow 128. An oil separator 130 outputs an oily solids flow 132 and the oil product flow 122. A solvent extractor 134 outputs a solids flow 136 and an oil/solvents mixture 138. Any heat that can be removed from mixture 138 is returned in a heated solvents flow 140. A last solvent recovery stage 142 outputs a spent shale flow 144 and is assisted with a waste water wash. A filter cake flow 148 is added to the last solvent recovery stage 142 with a wash water flow 150. Such filter cake flow 148 is produced by an extracted-oil solids filter 152 that removes solids from the oil product flow 120 and forwards filtered product oil flow 153.

The final solvent recovery from the aqueous leach system preferably includes a staged pressure relief let-down to atmospheric from around 500 psig

A distillation column 154 outputs a final-product synthetic crude oil 156. A heat exchanger 158 outputs a flow 160 and receives a flow 162. A flow 164 is forwarded for hydrogen production and sulphur recovery. A recycle solvent 166 is provided by the distillation column 154 to the pulverized oil shale ores flow 106. An H-donor flow 167 is added to flow 166 from an H-donor generator 168. A mid-distillate flow 169 from the distillation column 154 is sent to the H-donor generator 168. A hydrogen sulfide (H₂S) flow 170 is sent to a hydrogen plant sulfur recovery unit 171. Such produces a hydrogen flow 172, a pair of sulfur-free fuel-gas flows 173 and 178, and a sulfur flow 180. A fuel flow 173 is provided to run a power plant 174. One output from the power plant 174 is electricity 176. A sulphur flow 180 is output from the sulphur plant 171.

Guo Shu-Cai, et al., described some conversion steps like those of the oil shale ores processing plant 100 for an experiment they conducted. Such experiment was reported in “Conversion of Chinese Oil shale ores to Liquid Products using Supercritical Extraction,” pp. 311-316, German Publication: Erdol und Kohle-Erdgas-Petrochemie vereinigt mit Brennstoff-Chemie, Bd. 39, Haft 7, Juli 1986. Guo Shu-Cai, et al., found that the supercritical extraction of oil shale ores with toluene can give up to twice the oil yield over conventional retorting. When an H-donor was added to the supercritical solvent, complete recovery of the oil shale ores kerogen and high liquid-product yields was possible. But their paper did not describe a practical or complete system that recovered the solvent from the synthetic crude oil. What was described was that with a toluene/donor-solvent mixture with a tetralin content of about twenty percent, the oil yield (extract and liquids) can be enhanced as high as 200% of the Fischer Assay.

Earlier, in 1980, John Pratzer, II, was issued U.S. Pat. No. 4,238,315, (Pratzer '315) which describes the recovery of oil from oil shale ores. Such patent is incorporated by reference and may be helpful to the reader in implementing embodiments of the present invention. Pratzer '315 describes the extraction of oil from oil shale ores with the aid of elevated temperatures and pressures so that solvents with tetralin may be employed for highly efficient oil shale ores processing. The reactor effluent is described as being filtered, and the resulting filter cake rinsed with toluene. Again, such did not describe a practical or complete system that recovered the solvent from the synthetic crude oil.

The recovery and recycling of solvents with H-donors is described by Marvin Green, et al., in U.S. Pat. No. 4,325,803, (Green '803) titled “Process For Hydrogenation-Extraction Of Organics Contained In Rock,” and issued Apr. 20, 1982. A hydrogen transfer agent in its liquid phase is used to separate kerogen from oil shale ores. The oil separation occurs in a reactor at elevated pressure and temperature that keeps the solvent liquid. A slurry flash releases pressure and an adiabatic flash vaporization of the organic materials occurs. A portion of the vaporized materials is recycled as hot recycle oil vapors. Green '803 is also incorporated herein by reference as it may be helpful to the reader in implementing embodiments of the present invention.

In embodiments of the process of the present invention, the H-donor generator 168 starts with a mid-distillate fraction of naphthalene (C₁₀H₈) with a molecular-weight-128, e.g., in flow 169, and chemically reacts this with hydrogen using a catalyst. Such produces H-donor products in flow 167 like decalin (C₁₀H₁₈) (molecular-weight-138) and tetralin (C₁₀H₁₂) (molecular-weight-132) with boiling points of 190-220° C. Up to eight percent by weight of hydrogen is available for such chemical reaction. Typical oil shale ores production requires about two percent by weight hydrogen for the benefits described. About twenty percent of the total solvent mix used in the extraction/conversion is preferably H-donor mid-distillate fraction.

Method embodiments of the present invention combine crushed oil shale ores with a mixture of toluene or other low-boiling-point range organic solvent, and tetralin/decalin or other mid-distillate. Such is then fed into a slurry mixer and heated by a solvent recovery from a spent shale. Any water in the oil shale ores is eliminated. The slurry is then pumped with a recycle product oil stream into an autoclave where moderate temperatures and elevated pressures are used to convert substantially kerogen to a hydro visbroken stable crude oil with some gas production. The oil product is then separated from the spent shale under similar temperature and pressure conditions.

After separation, some of the oil product is used for recycling. The rest is distilled for the solvent mix by a low-boiling-point fraction and a mid-distillate fraction. Such mid-distillate fraction is hydrotreated and recycled to make up for any hydrogen used up. The spent shale is washed by a counter-current with low-boiling-point solvent fractions under elevated, e.g., supercritical, pressure and temperature conditions. Such enables easy separation of the solids and continued conversion of the residual kerogen. The final residue of spent shale with clean low-boiling-point solvent at supercritical temperature/pressure conditions is let-down gradually in stages with almost all the low-boiling-point solvent evaporating for recovery and reuse with the final stage sprayed with water to release any remaining solvent through steam stripping. The oil product is filtered before storage and is eventually piped to a commercial market.

Such gas produced is rich in methane, and thus is easy employed in the hydrogen production needed for mid-distillate hydrotreating. The hydrogen is used for heating and powering the rest of the process.

Some embodiments of the present invention include the use of at least one autoclave wherein high pressure leaching is used to convert kerogen to oil. The autoclave preferably includes an internal venturi draft tube to keep the slurry mixed. A pressurized extraction vessel continues the conversion process and acts to solubilize the converted oil. A series of pressurized solvent washing shale decanters are used in which shale moves counter-current to the solvent. Distillation columns, settling tanks and a plurality of pumps and heat exchangers are used to transfer and recycle components.

Referring again to FIG. 1, the plant 100 begins with the oil shale ores input feed 102 that preferably ranges in size up to forty inches. Such is forwarded to conventional commercial size reduction equipment to obtain a resulting ore of about ⅜-inch screen mesh. The crushed shale is fed by flow 106 into a slurry mixer system water remover 108, e.g., as described in Rendall '267, and incorporated herein by reference. The system water remover 108 receives heated solvent vapors from the solvent recovery stage 142 and the heat received helps maintain an operating temperature near the boiling point of water. Any water vapor is condensed and drawn off in waste water flow 110 which is flushed out in spent shale flow 144. The slurry is heated by heat exchanger 116 and the pressure is increased to 600 psig.

More heat can be added to flow 118 to raise the temperature of the flow 118 entering the kerogen converter 124 to about 400° C. The kerogen converter 124 preferably includes an autoclave and provides residence times of five to thirty minutes, depending on the oil shale ores ore. The kerogen is converted by pyrolysis with an H-donor distillate providing hydrogen to deal with olefin formation and unsaturated hydrocarbons. Some sulphur will detach as hydrogen sulfide.

The reaction is chemical and continues through the entire section under elevated temperature and pressure, e.g., kerogen converter 124, oil separator 130, and three-stage solvent extractor 134. The gas flow 128 produced by the reaction comprises methane, ethane, hydrogen, and some hydrogen sulfide. The oil flow 126 is separated in oil separator 130 by a pressure vessel.

The three-stage solvent extractor 134 can be implemented similar to the three pressure decanters 712, 730, and 740, in FIG. 2 of Rendall '267.

The oil is removed in flow 122. The oil separator 130 agglomerates the fines which settle out with the solids output flow 132. The hot product oil is fed to distillation column 154 via a heat exchanger 116 to heat the incoming slurry then through an extracted-oil solids filter 152 before the pressure is let down to provide the energy needed for distillation. The solids with oil exit in flow 132 to a three-stage solvent extractor 134, wherein fresh hot solvent at about 400° C. and about 700 psig is fed from flow 140 counter-current to the output spent shale.

Most of the remaining product oil output with the solvent in flow 138 via a heat exchanger 158 (for the incoming solvent) to the distillation column 154. The solids residue with some solvent leaves three-stage solvent extractor 134 via flow 136 to a solvent recovery solvent-residue solvent recovery unit 142. The solvent-residue solvent recovery unit 142 consists essentially of depressurizing the residue in a vessel fed from flow 136 thereby releasing most of the low-boiling-point solvent via flow 112 as a vapor to heat the incoming slurry at water removal water remover 108. The residue of spent shale is further cooled from about 200° C.-300° C. in a rotary drum with the remaining solvent from flow 148 joining flow 112. Filter residue from extracted-oil solids filter 152 feeds the depressurizing vessel in solvent-residue solvent recovery unit 142. The water from hydrogen plant 171 and/or waste water from water remover 108 can be used to cool the solids and dampen the spent shale residue for dust control during mine backfill. The solvent-residue solvent recovery unit 142 is any such system as described in Rendall '267 including depressurizing vessels and a cooling (rotary drum). The heat from the flow 136 is transferred via solvent and water (steam) vapors to the water removal water remover 108. It would aid water disposal to use acid water from the hydrogen/sulfur plant hydrogen plant 171 to cool the hot spent shale while waste water flow 110 is disposed of in flow 144 for dust control.

The recycle solvent to three-stage solvent extractor 134 flow 140 is fed from the distillation column 154 via flow 162 to a heat exchanger 158. Such is necessary at elevated pressure of about 400 psig. The heat can be provided by the oil/solvent output three-stage solvent extractor 134 via flow 138 at about 400° C. and leaves the heat exchanger 158 at flow 160 with temperatures about 150° C. to the distillation column 154. Auxiliary heat to flow 140 can also be provided by flow heat fuel gas from power plant 174. The hot oil from kerogen conversion from oil separator 130 flows at about 600 psig and 400° C. Flow 122 heats incoming slurry in heat exchanger 116 from which it leaves at about 600 psig and 150° C. An oil product flow 120 is forwarded via an extracted-oil solids filter 152. The filter residue is fed via flow 148 to the solvent recovery solvent-residue solvent recovery unit 142, for disposal of the fines. The filter can be of the metal porous cartridge type such as are readily commercially available, a pressurized rotary drum filter with engineered fabric for high temperatures about 150° C. such as supplied by CJ (Zyex Hi tech yarn) or any other suitable for process conditions. All are used in refineries for removal of catalyst fines before oil product distillation. The preferred route is a metal porous cartridge type. The distillation column 154 is fed the product oil flow 153 and recycle solvent at about 150° C. Flow 153 is a depressurized flow and additional heat can be provided by a fuel gas flow 178.

The distillation column 154 is conventional system with a mid-distillate metered off-take at about 200° C. in flow 169 for hydrotreating in H-donor 168. The recycle solvent off take at about 120° C. is metered to flow 166 feeding the slurry of the incoming raw oil shale ores for water removal in water remover 108. The ratio of mid-distillate flow 169 hydrotreated in H-donor 168 mixing with the recycle solvent flow 166 via flow 167 is about twenty percent of mid-distillate H-donor in the flow 114 proceeding under elevated temperature about 400° C. and pressure about 600 psig to the kerogen converter 124. The H-donor 168 for hydrotreating the mid-distillate from flow 169 is practiced by industry today, as referenced books for catalysts including, “Oil and Gas Journal Refining-Catalyst Compilations”. Such catalysts usually use alumina support with combinations of cobalt molybdenum nickel etc. as active agents. The technologies are similar to those described in “Petroleum Processing Handbook” edited by John S. Meketta, published by Marcel Dekken, June 1992 or “Upgrading Petroleum Residue and Heavy Oils” by Murray S. Greg, published by Marcel Dekker Inc., NY, N.Y. 1994. The hydrogen plant 171 receives hydrogen, hydrocarbon gases, including some light ends, some ammonia (NH₄) and hydrogen sulfide H₂S from flow 170.

The hydrogen is separated, concentrated, and reused by compressors. The hydrogen sulfide is converted to sulfur by a conventional Claus plant. Some of the fuel gas is used for further hydrogen production via a reformer for methane/light ends, and the rest for heat needed for process and electric power. About six to fifteen percent by weight of the kerogen is converted gases dependant on the source of the oil shale ores and the processing conditions for extraction of the kerogen. About two percent by weight of the produced oil is the hydrogen necessary for chemical reactions to produce a suitable pipelineable oil with the required viscosity and stability. Up to four percent by weight hydrogen has also been reported on particularly aromatic kerogens, producing more gas. The whole hydrogen system produces electric power. All these items in power plant 174 via flow 173 can only be quantified in specificity depending on the oil shale ores source and the size of the facility by those skilled in the art.

The Fischer Assay was developed for the oil shale ores industry to determine the efficiency of oil extraction processes. The Fischer Assay measures the recovery ratio of hydrocarbons from the oil shale ores. In prior art retorting processes, a Fischer Assay recovery of 80-100% is typical, and recoveries exceeding 100% are difficult to achieve.

Rendall '267, describes fifteen batch runs in which oil shale ores was treated with toluene under supercritical conditions ranging up to 400° C. and 1200 psig pressure for recovery of up to 120% of Fischer Assay of hydrocarbons.

In the batch runs, oil shale ores has been slurried in toluene in a batch stirred reactor and heated to temperatures up to 400° C. and held for periods ranging from zero minutes up to two hours.

Other data referenced on “carbonaceous” oil shale ores in which an H-donor mid-distillate has been used with the supercritical toluene shows kerogen conversion almost complete (about 95%). Such data clearly shows oil production with negligible olefins and hydrogen consumption from the H-donor mid-distillate at two to three percent by weight of the produced oil.

Method embodiments of the present invention for “carbonaceous” oil shale ores allow almost all of the kerogen to be produced as oil and gas (a 5-15% fraction). For example, Julia Creek, Queensland, Australia, oil shale ores, CRA report 1967-1988, show an average of 17-18% kerogen. The Fischer Assay yield is about H-donor 70 liters/ton, 14-15 gallons, representing about seven percent kerogen which is only 30% of what potentially could be available from the recovery method of the present invention. Such is adequately borne out by independent research on other Chinese, Australian and Eastern US shales shown in the references. The decrepitation observed on Colorado oil shale ores does not appear to occur in carbonaceous oil shale ores. However, continuous operation may change this phenomenon.

Some process embodiments of the present invention produce pipelineable synthetic crude oil from oil shale ores. A low-boiling-point organic solvent fraction is combined with an H-donating mid-distillate fraction. The temperature is raised to 370° C.-420° C. to make kerogen in oil shale ores soluble, and pressure conditions are raised to keep the solvents in their liquid phase or appropriate critical phase density at those temperatures. The solvent is recovered from the extracted kerogen in a solvent-recovery unit that is followed by a flash recovery that uses a pressure letdown and draws off the resulting solvent vapors.

It is important to use a one-quarter inch feed in cases of kerogen removal which do not powder or otherwise decripitate the host rock, e.g., as with Mahogany zone oil shale ores. Slurry pumpability under supercritical conditions will dictate. The process described in U.S. Pat. No. 4,737,267, can be used without the H-donor system.

A preferred embodiment to use the pressure and temperature of slurry residue after solvent extraction involves the final solvent recovery from aqueous leach system with staged pressure relief down to atmospheric from around 500 psig.

A pressure let-down efficiently recovers solvent and simultaneously leaches out mineral values with aqueous recycles. So most of the heat from the system can be recovered this way. The dry and damp residues can be used with the aqueous recycle for leaching out the mineral values.

A second preferred embodiment does not use solvents to wash the oil mixture from the residue, and instead uses high pressure, temperature hot water to separate the oil mixture from the residue.

The pressure let-down of the aqueous recycle efficiently recycles heat and residual solvent from the system. The residue is available for Portland cement production after washing in a pressure filter. The pressure and temperature reduction of the residue will release most of the water remaining in the solids.

The process continues in the removal of the washed residue by any suitable system such as pressure filters. The filtered leach liquor now contains the alumina and soda ash values ready for precipitation of the sodium carbonate monohydrate and of the aluminum trihydrate as is practiced today in industry both for specialty grade aluminum hydrates or metallurgical grade alumina for smelters.

The preferred requirement of the aqueous leach circuit leaving the pressure leach section of the residue is that is contains a saturated solution of the alumina and soda ash values, adjusted by water dilution of the recycle spent liquor. E.g., after primary removal of soda ash and alumina values). The recovery of soda ash values by crystallization as a mono hydrate is preferable at about 150° C. as is practices industrially today. (ref. Isonex, Jun. 11, 1997) The alumina values are then recovered with an alumina to sodium carbonate ratio of about 0.7 and an alumina concentration around 165 grams-per-liter as is practices in the Bayer process today (in notes of Don Donaldson Mar. 11, 1997). Alternatively the alumina (as aluminum trihydrate) is precipitated by sparging with CO₂.

The filtration particle size is critical. The industry has developed the art of precipitation control using recycle of fine aluminum hydrates as a seeding mechanism to control yield and size. Experimentation shows that the silica content of the produced alumina meets aluminum smelter industry requirements. Leaching with sodium carbonate, as opposed to caustic, will also reduce the aluminum sodium silicate formation. Free sodium hydroxide levels in leach liquor can also be controlled by converting to the carbonate with CO₂ sparging.

The chemistry involved in embodiments of the present invention is represented in the following Table.

B. Residue including alumina and soda ash formed at around 400° C.
as follows:
C. Residue including alumina and soda ash leached at temperatures
between 125° C. and 200° C. but typically about 150° C. with sodium
carbonate (or CO2) provides leach liquor at 150° C. Wherein
Na2CO3H2O, the monohydrate, is removed at concentrations around
thirty percent by evaporative crystallization at temperatures
around 100° C.
D. The leach liquor at around 70° C. including alumina: sodium
carbonate ratio of around 0.7 and with an alumina content of about
165 grams-per-liter can be reduced to about 80 grams-per-liter for
recycle by seeding with fine hydrate crystals (about 50% @ 44u or
825 mesh) at about 30 grams-per-liter, This removes half of
original alumina in the leach liquor, and then the liquor is
recycled for further leaching to replenish the alumina soda ash
values.
E. The residue in limestone-predominant kerogenous ore bodies has
unconverted carbon that is then calcined at over 1000° C. with
requisite amount of limestone to yield a Portland cement
equivalent. About 5% of the original kerogen will be unconverted
carbon. The calcining results in a clinker that can be ground to
the consistency of dry Portland cement.

FIG. 2 represents a process plants and method embodiment of the present invention for extracting multiple mineral products from oil shale ores, and is referred to herein by the general reference number 200.
Process 200 extracts crude oil from oil shale ores using supercritical extraction temperatures 370-420° C., with a preferred temperature of 400° C. The overall residence time at such temperatures is preferably 15-60 minutes. Process 200 further extracts soda ash from any dawsonite and nacholite associated with oil shale ores, and aluminum hydroxide clumps in various sizes. Bayer-type alumina and acid alumina can also be obtained from the hydroxides obtained in process 200. Aluminum metal is obtainable by reductive smelting of such alumina.

Process 200 begins with an oil shale ores feed material 202 that naturally and typically includes kerogen, dawsonite, nacholite, alumina and soda ash. Such oil shale ores is mined and transferred to a conventional particle-size reducer 204. The reduction is preferred to be less than one-quarter inch for easy slurry processing. A crushed oil shale ores 206 enters a supercritical solvent extractor 208, e.g., an extractor as described in U.S. Pat. No. 4,737,267. A hot residue 210 is then transferred to a pressure leach 212. Such receives a recycle 213 of leach liquor. Preferably, a temperature of 150° C. is used in pressure leach 212 with a residence time of about half hour, or at least in the range of 15-60 minutes. An output flow 214 is sent to a wash and pressure filter 216.

A washed filter cake output 218 can be returned to the oil shale ores mine for back fill. Such washed filter cake output 218 may instead be mixed and crushed with limestone in proportionate amounts to produce a material equivalent to dry Portland cement. A good commercial market exists for such materials. Conventional Portland cements are a type of hydraulic cement usually made by burning a mixture of clay and limestone in a kiln.

A leach liquor 220 is forwarded that includes alumina and soda ash values. A soda ash crystallizer 222 returns a liquor recycle flow 224 through a recycle storage 226. The soda ash crystallizer 222 also separates out a soda ash slurry 228 which is sent to a dryer 230. The concentration values are preferred to be greater than 300 grams-per-liter at 150° C. A typical dryer 230 uses evaporative crystallization at around 100° C., to yield sodium carbonate monohydrate crystals. Such crystals are centrifuged, washed, and dried to produce sodium carbonate, e.g., soda ash. A soda ash product 232 is output from process 200.

The concentration of the recycled saturated leach liquor is adjusted with a fresh water input 227, and is made dependent on the soda ash content of the residue available for leaching.

Such sodium carbonate monohydrate Na₂CO₃H₂O removal in soda ash crystallizer 222 is preferably limited to that necessary for a 0.7 ratio of alumina to sodium carbonate in the liquor. A liquor 234 at around 70° C. is then transferred to a precipitator-filter-washer 236. The feed preferably has concentration of about 30 grams-per-liter alumina, e.g., as aluminum trihydrate fines with 50% less than 44-microns. The concentration of alumina increases from 165-grams-per-liter to 195-grams-per-liter of the aluminum for precipitation at 70° C. This leaves a aluminum hydroxide residue with a concentration of about 80-grams-per-liter.

In alternative embodiments of the present invention, CO₂ is introduced under pressure at temperatures above 100° C. after the sodium carbonate monohydrate Na₂CO₃H₂O is removed at around 100° C. to 200° C. Such precipitates aluminum hydroxide. Crystalline granular aluminum hydroxide will be formed as long as the sodium carbonate values are about equal to the alumina values in the leach liquor. Increasing the residence time up to three hours increases the size of the fine particles in the precipitate. A further variation for use at atmospheric pressures includes seeding with fine aluminum trihydrate.

An agglomerate drier 238 can be manipulated to produce larger crystals by increasing the residence time up to 24-hours, e.g., as is common in the Bayer process. An output 240 produces dried aluminum trihydrate crystals. Such crystals can be commercially marketed for use as fire retardants, catalysts, and paper-making materials.

A flow 242 of trihydrate Al₂O₃3H₂O is both output from process 200 and input to a basic aluminum sulphate (BAS) converter 244. A BAS output in flow 246 is forwarded to a calciner 248. Fine crystals from agglomerate drier 238 can also be converted to BAS using sulphuric acid. Such acid is recycled in a flow 250. The BAS is calcined into an alumina and SO₂/SO₃ flow 252. The alumina is forwarded for a conversion 254 to aluminum metal, e.g., smelting. An aluminum metal 256 is output as a commercial product.

The supercritical solvent extractor 208 uses heat around 400° C. and pressures of around 500-psig to remove the organic carbon kerogen and breaks it down to a lighter oil 209. Solvent is recycled after hydrogenation, and a gas product 258 is sent to a utility 260 to produce a hydrogen flow 262 needed. A carbon dioxide gas (CO₂) flow 264 is sent to precipitator-filter-washer 236.

FIG. 3 represents an aqueous mineral extraction process embodiment of the present invention, and is referred to herein by the general reference numeral 300. The process 300 repeats much of what is illustrated in FIG. 1 for process 100. So the 100-series of reference numerals is repeated here for identical components. The differences are concerned with the flows in and out of oil separator 130 and the three-stage solvent extractor 134.

In process 300, the oil separator 130 has hot water 150 injected in remove oil from the solid residue. A barrier forms to float the oil out in flow 122 to heat exchanger 116 for heat removal. Such water causes light-ends in the oil within separator 130 to azeotrope. These are output in flow 112. This causes heat to be transfer to the incoming oil shale ores slurry 106 in water remover 108.

The residue from oil separator 130 is an aqueous slurry flow 132 forwarded to solvent extractor 134. Here it sits for the staged residence times necessary to leach out mineral values. The appropriate residence time is determined empirically, as it varies from one ore body to the next. In the case of Green River oil shale ores, one-half to one hour is adequate. The entire slurry can be sent for filtering in flow 136. Otherwise, the heat exchanger 158 of FIG. 1 can be used to conserve the heat in the aqueous supplement liquor. In any case, the liquor and the residence need filtration to provide clean liquor and a cake reside.

In general, oil shale ores that include kerogen and with other minerals associated with alumina and soda ash, are mined and crushed. This provides a slurry feed with solvents to an extraction module operated under supercritical conditions of pressure and temperature for the solvent. Such extracts the organic carbon content as oil. The residue is then leached with an aqueous sodium carbonate solution to remove the alumina and soda ash values. The remainder is washed and returned to the mine as backfill. The leach liquor is processed at around 100° C. to precipitate the soda ash values. The alumina values are recycled and replenished with residue remaining after oil extraction. The soda ash is dried to remove waters of hydration and then sent to market. The aluminum hydrate is then converted to alumina via conversion to a basic aluminum sulphate (BAS). Such is precipitated at around 200° C. at pressure. The BAS is calcined at around 850° C. into alumina. Sulphur oxide gases are recycled from a conversion to sulphuric acid.

In alternative method embodiments of the present invention, the crushed particle size of the oil shale ores is preferably one-eighth inch to one-quarter inch in size. The solvent used is a low-boiling type in the range of 100° C. to 140° C. An additional recycle of a hydrogen donor mid-distillate can be used such as tetralin, e.g., from a cut of the produced oil. The temperature is preferably maintained at 370-420° C., with an overall residence time of 15-60 minutes. But typically the residence time is about thirty minutes.

An aqueous recycle of sodium carbonate is preferably maintained at a temperature of greater than 150° C. at around 200 psig for the leaching of the residue. Sodium carbonate is leached into the recycle aqueous sodium carbonate leach liquor from the residue after oil extraction. Sodium carbonate monohydrate is preferably removed by precipitation using evaporative crystallization at temperatures about 100° C. After removal of the sodium carbonate, the leach liquor is sparged with CO₂ to precipitate aluminum trihydrate crystals. The amount of sparging CO₂ used is about equal to that entering from the oil shale ores residue. Such sodium carbonate monohydrate is preferably dried with the production of dense soda ash.

The aluminum trihydrate crystals are washed, dried and calcined into alumina at about 850° C. to 950° C. Fine aluminum trihydrate is recycled to for the production of crystalline aluminum trihydrate from the leach liquor. An alternative embodiment precipitates the aluminum trihydrate using fine aluminum trihydrate (1) at a 1:4 ratio, (2) a temperature around 65° C., (3) a concentration of alumina of about 160 grams per liter, and (4) a ratio of alumina to sodium carbonate of about 0.7.

The precipitated aluminum trihydrate is preferably agglomerated into larger crystals using starch with a residence time in thickness of about 20-25 hours. The aluminum trihydrate production is preferably directly converted to basic aluminum sulfate using a recycle acid stream from the calciner off-gasses at a temperature of around 200° C. with a corresponding pressure for aqueous systems. The recycle leach liquor is preferably purged after any depletion of soda ash and alumina values to remove accumulation of impurities in the aqueous leach circuit.

Process embodiments of the present invention extract hydrocarbon products, alumina, and soda ash, from oil shale ores. Such includes extracting various amounts of such chemicals, all based on supercritical solvent extraction of most of the hydrocarbons at temperatures around 400° C. The alumina and soda ash values are leached out with aqueous sodium carbonate leaching at reduced temperatures of around 150° C. with a corresponding reduced pressure. The soda ash monohydrate values are precipitated from the leach liquor at around (over) 100° C. with the alumina values precipitated using CO₂ at lower temperatures (or the same) and alternatively using recycled fine aluminum trihydrate at 65° C. under conditions practiced today to produce alumina. Alternatively and especially should silica prove to be a high contaminant the precipitate aluminum hydroxide is preferably converted to acid alumina by an acid recycle stream dissolving the alumina (leaving the silica to be filtered out) and precipitating basic aluminum sulfate at about 200° C. (250 psig) for calcination at around 900° C. to produce alumina and recycle of the sulphate oxide gasses via a sulphuric acid plant.

To summarize, the solvent conversion of kerogen in the oil shale ores to crude oil using high pressure around 500 psia with temperatures around 400° C. and its separation from the host rock by solvent extraction or aqueous flotation is described in FIGS. 1 and 3. The solvent/mid distillate fraction of the oil extracted is recycled, and the gas production is cleaned of ammonia and sulphur in 171 before it is used for hydrogen needs of the process embodiments and for the power plant means 17.

The process embodiment FIG. 2 removes the minerals from the residue by high pressure/high temperature (150° C.-300° C.) aqueous means containing soda ash to access valuable minerals in Green River oil shale ores (alumina and soda ash), Julia Creek oil shale ores (alumina and molybdenum), Estonia oil shale ores (alumina) with a residue calcined for production of Portland cement.

The processes in FIGS. 1-3 use the recycle solvent to carry the hydrogen-donor mid-distillate in a slurry and to convert the kerogen in the oil shale ores into pipelineable oil products. The water in the oil shale ores is removed prior to conversion of the kerogen to oil. The recycle solvent is entirely vaporized in the distillation column and then condensed for recycle to slurry the oil shale ores. The process in FIGS. 4-5 does not preliminarily remove the water in the oil shale ores, but rather adds more to the process. This has the beneficial effect of reducing the need for hydrogen.

FIG. 4 represents the physical implementation of a kerogen extraction method embodiment of the present invention. An oil shale ores plant 400 embodiment of the present invention inputs oil shale ores ore with its intrinsic water in a flow 402 to a size reducer 404. A +245° C. recycle 108 from a holding tank is added to the input oil shale ores in a material conditioner 410. A heat exchanger (H/E) 412 takes heat away from a flow 414, as does a H/E 416 from an oil flow 418. An H-donor flow 419 is added to this heated and pressurized oil shale ores. The salvaged heat is put to use in a reactor-furnace 420 which raises the temperature of the slurry to 450° C. A separator 422 produces a hot oil flow 424 a residue flow 426. A pulsed-wash column 428 takes one more shot at removing the oil from the solid residues. Any extracted oil is directed in a flow 430 where H/E 416 recaptures their heat. A recycle solvent 432 is also used by the pulsed column 428. The remaining solid residue is relieved of its processing heat by H/E 434 before slurry depressurizer 436. A recovered solvent flow 438 is returned to a distillation column. The remaining solid residue again relieved of pressure by a depressurizer 440. Another recovered solvent flow 442 is also returned to the distillation column. A last pressure step down stage 444 recovers a solvent flow 446 that is returned to the distillation column in a combined vaporized solvent flow 448. Alternatively, the solvent flow 448 is condensed in cooler 482.

The dry solid residue is output in a flow 450. The oil extracted in flows 414 and 418 are let-down by a depressurizer 452 for a distillation column 454. A solvent in a flow 455 is recovered from separator 422 and let down by another depressurizer 456 for distillation column 454. A sludge flow 457 from distillation column 454 is sent to the pulse-wash column 428 to recover product oil before disposing of the solid residue 450.

A gas flow 458 is produced by distillation column 454. A separator 464 removes oil output in a flow 462 from gas after cooling in 460 and gas exits in a flow 468. Oil/solvent returns to the distillation column 454 in flow 466. Any water is removed and output in a flow 470. A sulphur removal stage 472 cleans up a gas output in a flow 474 for use by utilities or a hydrogen-plant 476.

The bottom of the distillation column 454 produces a +245° C. flow 496 for recycle in the process. A filter 497 is used to extract product oil output in a flow 498.

The principal product produced by plant 400 is a high quality synthetic crude oil. Compared to other processes, the oil product of plant has greatly reduced levels of sulphur, nitrogen, and oxygen.

In general, the oil shale ores input feed 402 is crushed by size reducer 404 before being fed to slurry conditioning vessel 410. It stays there long enough to make a slurry. Such includes most of the intrinsic water in the raw oil shale ores. Other competing processes remove this water. The slurry receives flow 408 from a +245° C. recycle tank (not shown) and +245° C. recycle flow 496. The mid-distillate donor flow 493 is hydrogenated by standard hydrotreating, e.g., using hydrogen plant 476 to produce hydrogen for hydrogenation of the mid-distillate flow 493 and flow 419. Such H-donor flow 419 could be injected anywhere between slurry conditioner 410 and gas-fired reactor-furnace 420, inclusive.

Solvent vapor flow 479 is sent to a fractional distillation separator for the desired 110° C. fraction for solvent recycle flow 484. The rest of the solvent 486 is sent to the product oil storage. The cooler 482 provides the necessary solvent liquid for both separator 480 and flow 478 to the distillation column.

The mid-distillate vapors in flow 489 going into the fractional distillation separator 490 provide the 180-240° C. mid-distillate fraction flow 493 used in hydrogenation and for recycle flow 419 to the process. The excess of this mid-distillate is skimmed off as product oil for storage in flow 494. Cooler 492 condenses the necessary mid-distillate liquids for separator 490, the remainder return as flow 488 to the distillation column 454.

Other means can be used as substitutes for separators 480 and 490, e.g., by drawing liquids instead of vapors off the distillation column and then using a separating heat of steam, etc.

The slurry is heated by heat recovered by H/E 412 and 416, and by cooking in the gas-fired reactor-furnace 420, to the kerogen conversion temperature, e.g., 842° F. (450° C.). Separator 422 could be implemented with conventional individual gas and liquid separators.

The solid residues are washed free of oil by solvent from flow 484 and flow 432 in the pulsed-wash column 428 heated to 847° F. (450° C.) and at supercritical pressure conditions, 600-650 psig. Such heat comes from the heat salvaged by H/E 434 and supplemental gas-fired furnaces.

The slurry solid residue is depressurized from as much as 650 psig, e.g., in the three let-down stages 436, 440, and 444. Each pressure let-down vaporizes more solvent, e.g., flows 438, 442, and 446. The dry solid residue 450 is thus solvent free.

The hot, vaporized solvent flow 448 may be connected directly to distillation column 454, as shown in FIG. 4, or alternatively to H/E cooler 482 to recycle this fraction, e.g., directly in flow 484, or otherwise.

Depressurizers 452 and 456 accommodate the pressure changes needed by distillation column 454 for oil flows 414, 418, and solvent flow 426. The high pressure/high temperature zone surrounding reactor-furnace 420 operates at 650 psig and 847° F. (450° C.)

The distillation column 454 operates at 30-250 psig, and preferably around 50-psig. Under such conditions, the +245° C. fraction will exist as a liquid with the lighter fractions fractionated to provide a mid-distillate 494. Such has a boiling-point (BP) of 180° C. to 210° C. and is used for hydrogenation. The amount needed is about 150% of the kerogen content weight of the oil shale ores. Typically, kerogen amounts to 18-20% of the oil shale ores. The lighter solvent fraction (BP 110° C. to 170° C.) is fractionated for recycle flow 484. The fractionation for particular boiling-point materials for mid-distillate comes from flow 489 through separation column 490 via cooling H/E 492 to provide for fraction flow 494.

The fractionation for the solvent recycle 484 is provided from flow 479 and separated in column 480 with cooling provided by H/E cooler 482. The product oil is the gain in volume of solvent 484, mid-distillate 494, and +245 C recycle flow 496. The product gas flow 426 from separator 422 is depressurized in unit 456 and fed through the distillation column 454 to remove solvent. It is then cooled in flow 458 by cooler 460 to remove even more solvent. Condensed water is removed in flow 470 by separator 464.

After sulphur removal stage 472, hydrogen can be made using conventional steam reforming in H-plant 476, and with steam/electric production.

In embodiments of the process of the present invention, the H-donor generator 494 starts with a mid-distillate fraction of naphthalene (C₁₀H₈) with a molecular-weight-128, e.g., in flow 419, and chemically reacts this with hydrogen using a catalyst. Such produces H-donor products in flow 167 like decalin (C₁₀H₁₈) (molecular-weight-138) and tetralin (C₁₀H₁₂) (molecular-weight-132) with boiling points of 190-220° C. Up to eight percent by weight of hydrogen is available for such chemical reaction. Typical oil shale ores production requires about two percent of the oil shale ores by weight of hydrogen for the benefits described.

The kerogen content of Colorado USA and Julia Creek Queensland Australia oil shales is typically 18-20%, and needs about one-and-a-half times its weight as H-donor to carryout the chemical reactions that will produce a high quality synthetic crude from the +245° C. fraction at 445° F.

Method embodiments of the present invention crush oil shale to a quarter-inch and combine it with a +245° C. fraction of the product oil to make a slurry. At start-up, 18-20 cycles of +245° C. fraction recirculation are needed to self-generate enough of the +245° C. fraction to balance the system. The slurry residence time in the mixer 410 is set at about three minutes. The short residence time and a dearth of lower boiling-point solvents rules out water removal by azeotropic action. The slurry is pumped at a pressure balance of 650-700 psig though H/E's 412 and 416. The oil shale slurry proceeds to the furnace reactor-furnace 420 with the H-donor mid-distillate from flow 419. Up to 10% of the oil shale in extra water is added from flow 419. The slurry leave the reactor-furnace 420 at about 650 psig and 450° C. for separation of the gas product in separator 422. The gas leaves as flow 426 and the liquid flows to H/E 412. Solids are sent to pulsed-wash column 428. The solid residue has the oil washed clear by supercritical low boiling-point solvent (110-120° C.) at 650 psig at 450° C. Mixer/settlers could be used instead of a pulsed-wash column (where the solids flow downward through an upwelling of supercritical solvent medium). The H/E 434 provides heat taken from the solvent flow.

Once the solvent in the spent shale slurry has dropped below the 300° C. supercritical temperature, it is discharged into a flash evaporator 436 which drops the pressure to about 200-psig. The next stage 440, uses a rotary drum and drops the pressure to 75-psig. A final rotary drum stage 444 drops the pressure down to 10-psig and flashes the remaining solvent into vapors. All the solvent vapors in flow 448 are condensed for recycle.

The +245° C. recycle oil and product oil is removed from the liquid-solids separator 422 in flow 424 and is let-down to 100-psig by depressurizer 452 before entering distillation column 454 at about 50-psig. The solvent washed from the solids in flow 430 is also let-down to 100-psig by depressurizer 452 before entering distillation column 454 at about 100-psig. The +245° C. fraction remains liquid in the distillation column 454 and is recycled to slurry the incoming shale ore in flow 402. The fines that accumulate in the distillation column 454 as sludge are removed in a flow 457 and then cleaned in pulsed-wash column 428 to recover oil. The fines are eventually disposed of in dry solid residue flow 450.

The oil product is a combination of the −245° C. fractions which were hydrotreated-hydrocracked to a quality low in nitrogen (N₂), sulphur (S₂), and oxygen (O₂), and some +245° C. fraction. It is presumed a catalytic reaction occurs as a result of the oil shale clays that are mixed in the H-donor mid-distillate and hydrogen produced by the water/steam reactions. Such oil shale clays usually include alumina, a common catalyst used in many reactions. The mechanism at work has not yet been identified exactly, but it has been verified as functioning this way in observations.

The recirculating +245° C. fractions contain most of the nitrogen (N₂), sulphur (S₂), and oxygen (O₂), are recirculated repeatedly into the hydrotreating-hydrocracking hydrogen donation regime of the reactor-furnace 420. Each cycle through halves the nitrogen (N₂), sulphur (S₂), and oxygen (O₂), and produces lighter products. The repeated recycle before exiting as product oil results in low gravity, high quality synthetic crude oil. Filter 497 cleans the last bits out to purify product oil output flow 498.

All of the kerogen made available is effectively converted into product oil because the furnace reactor and separator are allowed thirty minutes to operate on it at 450° C. and 650 psig.

FIG. 5 refers to an experimental 1-ton test plant process 500 to illustrate the action going on in oil shale ores plant 400 (FIG. 4). A 0.5 barrel per hour feed of H-donor mid-distillate (H2) 502 is added to a recirculating carrier 504 of about nine barrels per hour of +245° C. mid-distillate. These correspond to H-donor flow 419 and +245° C. recycle flows 408 and 496, respectively, in FIG. 4. From the mining operations, a one-ton (1T) feed of oil shale ores 506 (oil shale ores 402 in FIG. 4) is mixed in. The 10.5 barrel per hour mix 508 will output about 0.75 ton of residue 510, e.g., dry solid residue 450 in FIG. 4. The useful products from the process 500 include 0.5 million cubic feet (MCF) gas 512, and one barrel per hour of product oil 514. These correspond to gas flow 474, and product oil output flows 486 and 498, in FIG. 4. To balance the flows, a mid-distillate flow 516 is taken off. Practically any size oil shale ores processing plant can be scaled up from process 500.

Tests that were recently conducted indicate that embodiments of the present invention are particularly efficient in the removal of product oils from carbonaceous oil shale ores with intrusions of cannel coal. It is speculated that the combination of a catalytic action of the clays in the oil shale, and the hydrogen donor mid-distillate raised to the 450° C. and 650 psig supercritical conditions, is responsible for the surprising yields that were measured to exceed 350% of the Fischer assays of two sources of samples.

Although the present invention has been described in terms of the presently preferred embodiments, it is to be understood that the disclosure is not to be interpreted as limiting. Various alterations and modifications will no doubt become apparent to those skilled in the art after having read the above disclosure. Accordingly, it is intended that the appended claims be interpreted as covering all alterations and modifications as fall within the true spirit and scope of the invention.

Claims

1. A method for converting ores into a product oil, comprising:

recirculating +245° C. fractions around through an oil shale slurry pressurized to 650 psig in a reactor heated to 450° C., depressurizing them for separation in a distillation column, and returning them with an H-donor mid-distillate.

2. The method of claim 1, further comprising:

not dewatering oil shale ores contributed to said oil shale slurry; and

adding water to said reactor.

3. The method of claim 1, further comprising:

operating said distillation column at about 50-psig.

4. The method of claim 1, further comprising:

using clays included in a feed of oil shale ores as a catalyst that substantially increases yields of product oil.

5. An ore processing plant, comprising:

an input for conditioning a mixture of incoming ore and a recirculating hydrogen-donor mid-distillate into a slurry;

a reactor operating under supercritical temperature and pressure conditions for converting kerogen in said ore into product oil, mid-distillates, solvents, and gas;

a distillation column connected to receive and separate said product oil, mid-distillates, solvents, and gas;

an output for product oil; and

a recycle for hydrogenation of said mid-distillates from the distillation column and for use in the input to make said slurry continuous.

6. The ore processing plant of claim 5, further comprising:

a separator included with the reactor for separating product oil from the gas and solid residues; and

a recycle of solvents from the distillation column for use in supercritical solvent extraction of the oil residue in a spent oil shale.

7. The ore processing plant of claim 6, further comprising:

a pulsed-wash column connected to receive solid residues and sludge from the separator and distillation column, and for using recirculating solvents to wash out product oil, and for removing and disposing of solid residues from the processing.

8. The ore processing plant of claim 5, further comprising:

means for adding water to the reactor;

wherein, water is not removed from raw oil shale ore at the input before being added to said slurry.

9. The ore processing plant of claim 5, further comprising:

means for using the equivalent of +245° C. fraction in the recycle.

10. A continuous conversion plant for producing product oil from mined ores, comprising:

an input for conditioning a mixture of at least one of incoming cannel coal, and oil shale ore, and a recirculating +245° C. fraction hydrogen-donor mid-distillate into a slurry;

a reactor operating under supercritical temperature and pressure conditions of 450° C. and 650 psig for converting kerogen in said ore into product oil, mid-distillates, solvents, and gas;

a distillation column connected to receive and separate said product oil, mid-distillates, solvents, and gas;

a separator included with the reactor for separating product oil from the gas and solid residues;

a recycle of solvents from the distillation column for use in supercritical solvent extraction of the oil residue in a spent oil shale.

an output for product oil; and

a recycle for hydrogenation of +245° C. fraction mid-distillates from the distillation column and for use in the input to make said slurry continuous.

11. The plant of claim 10, further comprising:

a pulsed-wash column connected to receive solid residues and sludge from the separator and distillation column, and for using recirculating solvents to wash out product oil, and for removing and disposing of solid residues from the processing; and

means for adding water to the reactor;

wherein, water is not removed from raw oil shale ore at the input before being added to said slurry.

12. The plant of claim 10, wherein, for each ton of ore being input, 0.5 barrel per hour feed of H-donor mid-distillate is added to a recirculating carrier of about eleven barrels per hour of +245° C. mid-distillate, a 12.5 barrel per hour mix outputs about 0.75 ton of residue, and the useful products from the process include about 0.5 million cubic feet (MCF) gas, and one barrel per hour of product oil, and to balance these flows, a mid-distillate flow is taken off.