AU657841B2 - Production of hydrogen - Google Patents

Description

657841 P/00/0011 Regulation 3.2

AUSTRALIA

Patents Act 1990 COMPLETE SPECIFICATION FOR A STANDARD PATENT

ORIGINAL

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Name of Applicants: Actual Inventor: Address for service in Australia: COMMONWEALTH SCIENTIFIC RESEARCH ORGANISATION Dr. Nguyen Vu Dung CARTER SMITH BEADLE 2 Railway Parade Camberwell Victoria 3124 Australia PRODUCTION OF HYDROGEN AND INDUSTRIAL Invention Title: Details of Associated Provisional Application: Australian Patent Application No. PL 1183 filed 5 March 1992 by COMMONWEALTH SCIENTIFIC AND INDUSTRIAL RESEARCH

ORGANISATION

The following statement is a full description of this invention, including the best method of performing it known to us h, p- -2- TITLE: PRODUCTION OF HYDROGEN The present invention relates to a method for producing hydrogen, to catalysts for use in the production of hydrogen and to a process for the production of synthetic crude oil from oil shale utilising hydrogen as produced using the process and catalyst of the invention.

Known reserves of oil shale are very large. For example, the amount of oil that could be extracted from oil shale reserves in tL, United States has been estimated to be equivalent to fifty times the known reserves of oil in the United States. Large deposits of oil shale are also found in Australia.

The production of synthetic crude oil from oil shale involves a number of processing steps. Figure 1 shows a conventional recycle shale process for the production of oil from oil shale. Following mining and preparation of the shale (omitted from the Figure for clarity), the shale is dried in drying step 10 and sent to retort 12. The shale is heated in the retort whereupon the organic component of the oil shale is largely vaporised. This vapour is cooled to condense the oil in oil recovery process 14. Following retorting, the spent shale is combusted in combustor 16 to provide hot solid (17) as the heat carrier for the retorting step. The excess combusted shale, after being passed through a heat exchanger 18 to recover heat, is then disposed of in an appropriate manner e.g. land fill.

The oil recovered from the retort off gas has a low H content and high N S and 0 contents and requires further processing before a useful product is obtained. Upgrading of the raw oil is carried out by the treatment of the raw oil with hydrogen. The processes involved in the oil upgrading are shown in 0 0 25 oil upgrading block 20 of Figure 1. The hydrogen required in the hydrotreating step is produced in hydrogen plant 22 by the steam reforming of the retort gas, oo supplemented with hydrotreated and natural gas. A simplified flow diagram of a conventional steam reforming process is shown in Figure 2. The process consists of four main steps: Sulphur removal Catalytic steam reforming (32); Water Gas shift and gn\02\12738.c 43 93 i- -3- Gas purification (36).

To protect the catalysts in the hydrogen plant, the hydrocarbons have to be desulphurised before being fed to the reformer. The desulphurised feed stock is then mixed with processed steam and reacted over a nickel based catalyst contained inside a system of high alloy steel tubes. The reforming reaction is strongly endothermic, with energy supplied by radiant combustion of fuel gas or oil. The metallurgy of the tubes usually limits the reaction temperature to 800-920'C. To obtain low concentration of hydrocarbons in the product gas, the process generally employs a high steam to carbon ratio low pressure (2.5 3.5 MPa) and high temperature (850-890'C).

After the reformer, the product gas mixture of carbon monoxide and hydrogen passes through a heat recovery step and is fed into a water gas shift converter (34) to produce additional hydrogen. The main by-product of the process, carbon dioxide has to be removed to give the desired hydrogen purity.

The residual carbon monoxide and carbon dioxide remaining in the hydrogen stream after carbon dioxide removal are converted to methane in methanotor 38.

i° The production of synthetic crude oil from oil shale using present conventional plant is very capital intensive. A 1983 study found that retorting o o' and upgrading of the raw oil to synthetic crude oil are the two major cost S" 20 contributing blocks.

The 1983 study of capital costs for upgrading the raw oil to synthetic Oo'O crude oil gave the following: of Capital Cost for Oil Upgrading Hydrotreating 46 Hydrogen production :W 25 Retort gas treating 9 Sulphur recovery 6 Sour water treating 4 100 Hydrotreating and hydrogen production are the two major contributors to the cost of upgrading the raw oil.

Although the capital costs for the production of hydrogen by conventional gn\02\12738.c 4 3 93 -4steam reforming of light hydrocarbons is high, the technology is mature and competitive. The steam reformers used in the process are very large, principally because radiant heat transfer is used to provide the heat of reaction, requiring large tube spacing. Cost savings could be achieved by replacing the costly fsteam reformer and its associated water gas shift converter and gas purification steps with a simpler, more efficient and therefore more economical system.

It is an object of the present invention to provide an alternative process for the production of hydrogen.

SThe present inventors have made the unexpected discovery that combusted oil shale has catalytic activity for the dehydrogenation of light hydrocarbons to carbon and hydrogen.

As used hereinafter, the term "oil shale" refers to inorganic material which is predominantly clay, shale or sandstone in conjunction with organic materials composed mainly of carbon, hydrogen, sulphur, nitrogen and oxygen.

The organic material found in the oil shale is called kerogen.

As used hereinafter, the term "spent shale" refers to oil shale which has been retorted to liberate volatizable hydrocarbons to leave an inorganic material containing residual organic matter composed of carbon, hydrogen, nitrogen, sulphur and oxygen. The residual organic matter is called char.

As used hereinafter, the term "combusted shale" refers to spent shale in which the char has been at least partly combusted.

According to a first aspect, the present invention provides a process for :the production of hydrogen by the dehydrogenation of light hydrocarbons comprising contacting the light hydrocarbons with combusted oil shale under conditions suitable for a dehydrogenation reaction to occur. The term "light hydrocarbons" is taken to encompass C 1

C

5 hydrocarbons.

In the process of the present invention, the light hydrocarbons are contacted with combusted oil shale and are converted to hydrogen gas and carbon. The carbon produced by the dehydrogenation of the light hydrocarbons may be laid down as a solid on the combusted oil shale.

gn\02\12738.c 4 3 93 Without wishing to be bound by theory, the present inventors believe that the combusted oil shale exhibits catalytic activity towards the dehydrogenation of the light hydrocarbons.

The dehydrogenation reaction may take place at a temperature of from 500 1,200°C, preferably 750 950°C, a pressure of from 0.1 atmospheres, preferably 1 20 atmospheres and a catalyst weight time of from 1 to 20000 kg of catalyst.hour per kmol of light hydrocarbons, preferably 2000 kg of catalyst.hour per kmol of light hydrocarbons, a catalyst residence time of from 0.5 120 minutes, preferably 3 30 minutes. The combusted shale may have a carbon content of from 0 to 20 wt preferably from 1.5 wt The light hydrocarbon feed is preferably selected from the group comprising retort gas, natural gas, hydrotreater gas and mixtures thereof. By retort gas, it is meant the gas stream leaving the oil recovery unit. In this unit, retort off vapour is separated into raw oil and retort gas. By hydrotreater gas, it is meant the gas stream exiting the hydrotreater unit.

According to a second aspect the present invention includes a catalyst for the catalytic dehydrogenation of the light hydrocarbons comprising combusted oil shale.

The process and catalyst of the first and second aspects of the invention are particularly suitable for use in the production of synthetic crude oil from oil S* shale.

Accordingly, in a third aspect the present invention includes a process for the production of synthetic crude oil from oil shale comprising: o supplying oil shale to a retort; (ii) heating said shale in said retort to produce retort off-vapour and spent shale; (iii) combusting said spent shale to produce heat and combusted shale; (iv) passing said retort off-vapour through an oil recovery stage to separate raw oil from retort gas; and upgrading said raw oil to synthetic crude oil by treating said raw gn\02\12738.c 43 93 -6oil with hydrogen in an hydrotreating process wherein said hydrogen is produced by the dehydrogenation of a light hydrocarbons-containing gas in the presence of combusted shale.

Preferably the light hydrocarbons-containing gas of the first, second and third aspects of the present invention comprises said retort gas or a mixture of said retort gas, natural gas and gas from said hydrotreating process.

The combustion of the spent shale to produce the combusted shale may be carried out under a wide range of combustion conditions suitable for at least partly combusting the char of the spent shale. For example, the combustion temperature may range from 500-1500'C, preferably 750-1000'C and the pressure may range from 0.1 50 atmospheres, preferably 1 20 atmospheres.

The inlet gas fed to the combustor may have an oxygen content ranging from 1% to 100%. The gas fed to the reactor may be pre-heated or it may enter at ambient temperature. Preferably, the oxygen content of the gas feed to the 15 combustor is low, with a range of 2.0 to 21.0% oxygen being especially o arser 000,, preferred. The residence time of the shale in the combustor may range from seconds to 2 hours, prefcrably 30 to 600 seconds.

0 Further aspects of the present invention will now be described with reference to the Examples and drawings in which: Figure 1 is a process flow sheet of a conventional oil shale 0 0 treatment plant; 0 Figure 2 is a flow sheet of a conventional hydrogen plant utilising °000 steam reforming; 00 Figure 3 is a flow sheet of a conventional oil shale treatment plant; and Figure 4 is a flow sheet of a modified oil shale treatment plant utilising hydrogen produced according to the present invention.

It will be appreciated that the Examples and drawings illustrate preferred embodiments of the invention and that the invention is not to be construed as gn\02\12738.c 4 3 93

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-7being limited by the Examples and drawings.

Example 1 A sample of spent shale was combusted in a 40 mm i.d. fluidised bed combustor to produce combusted shale. The combustion conditions were as follows:- Temperature, 'C Inlet 02 concentration Pressure, MPa Solid Residence Time, s 2.9 0.1 The combusted shale so produced had the following properties: Surface area, m 2 /g 32 Total carbon, 3.53 Inorganic carbon, 0.09 Total hydrogen r 13 Total sulphur, 1.30 Total nitrogen, 0.22 0.0 0000J 0000 o -3 300 The shale was obtained from the Stuart deposit located in the State of Queensland.

Experimental runs to determine the catalytic effect of the combusted oil shale on the dehydrogenation of light hydrocarbons were carried out in a conventional high pressure apparatus. The flow rate of methane was controlled by electronic niass flow controllers. Product analysis was performed by on-line gas chromatography (Shimadzu GC-8A, Hewlett Packard 3392A integrator) using a thermal conductivity detector and a carbosphere 80/100 column. Water was removed from the reaction products by a trap at the reactor exit.

The reactor used for this Example was constructed from a 7 mm o.d. by 4 mm i.d. quartz tube which was vertically mounted inside a high temperature furnace with a heated length of 150 mm. Two sets of quartz wool packing (one at the bottom of the reactor and the other at the top) were used to hold combusted shale within the heated region of the reactor.

The apparatus was used to carry out experiments on the dehydrogenation of methane to hydrogen and carbon in the presence of combusted oil shale. The gn\02\127 38 .c 4 3 93 I eperi l c n ad n-al -8experimental conditions and results are shown below: Temperature, °C Pressure, MPa Methane flow, cc/min STP Oil Shale Combusted shale amount, g Particle size, [tm Run time, min Methane converstion, 890 0.1 13 Stuart 1.0 100 140 17.7 890 0.1 13 Stuart 100 200 10.2 Example 2 A sample of spent shale was combusted in a 150 mm i.d. fluidised bed combustor at the following conditions to produce three samples of combusted shale.

0 -'00 0 00 050 00 00 a a 0 0o 0 ft o Condition Combustion Temperature Inlet 02 concentration (vol%) Pressure (MPa) Solid residence time (min.) 1 2 3 816 820 865 3.7 5.8 2.8 0.1 0.1 0.1 10.9 6.7 7.1 B ~d O(I Il u.l i i Jil The shale was obtained from the Rundle deposit located in the State of Queensland, Australia. The samples of the combusted shale so produced had the following properties: Sample 1 2 3 Combustion condition 1 2 3 Surface area (m 2 26.7 33.5 36.0 25 Total carbon 1.00 2.67 4.26 Inorganic carbon 0.03 0.06 0.07 Total hydrogen 0.01 0.04 0.07 Total nitrogen 0.09 0.16 0.23 Total sulphur 0.54 0.87 1.22 Total iron 6.2 6.1 gn\02\12738.c 4 3 93

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-9- In order to demonstrate the catalytic effect of combusted shale on the dehydrogenation of light hydrocarbons, a reactor was constructed from a 500 mm long straight quartz tube (13 mm A fixed bed of catalyst (combusted was supported by a section of quartz wool placed at the bottom of the reactor. Another section of quartz wool placed on top of the catalyst bed was used to provide preheating of feed gas and to reduce the reactor dead volume.

For some experiments, a thermocouple well (2 mm dia. Quartz tubing) could be located at the center of the reactor to allow measurements of temperature along the length of the bed. The reactor was mounted vertically inside a 1 kW three zones temperature controlled tube furnace (350 mm long x 26 mm The reactor with about 7 g of Rundle combusted shale having particle sizes between 0.57 and 3.00 mm Was initially heated to the required temperature under a continuous nitrogen gas flow of about 10 cm 3 /min. Feed gas was switched from nitrogen fo the desired feed gas as soon as the reactor temperature had 15 stabilized. The pressure was maintained at one atmosphere.

A small amount of the product gas was sampled by an on-line portable M200D gas chromatography (GC) unit at intervals of one minute for the duration of the experiment. Sampling of the product gas was started at the same time the feed gas was switched from nitrogen to methane. The 20 concentrations of H 2 02, N2, CH 4 and CO in the product gas were analysed using the GC's Mol Sieve 5A column. The HayeSEP column was used to analyse for the concentrations of CO 2 and ethane in the product gas.

The product gases were continuously collected in a sealed plastic bag whose expansion caused a continuous displacement of water. The weight of displaced water was continuously measured by an electronic balance (A&D FP- 6000) and logged by an IBM compatible computer to obtain the product gas volumetric flowrate.

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gn\02\12738.c 43 93 10 The experimental conditions and results are shown below: Run Sample Temperature Methane Methane Number Number 0 C) feed rate Conversion (cm 3 /min) 1 1 900 36.9 13.4 2 2 900 35.4 16.2 3 3 900 35.1 20.9 4 1 940 17.2 28.8 2 940 17.2 34.6 6 3 940 17.4 38.5 7 3 900 7.2 47.4 8 2 940 7.2 51.4 Example 3 Further experiments to demonstrate the catalytic dehydrogenation were carried out using a reactor constructed from a 1" Schedule 40 253MA stainless steel pipe. It was mounted vertically inside a five zone 3 kW temperature controlled tube furnace (0.9 m long x 70 mm Rundle combusted shale (Sample No. 2) moved as a packed ,ed from feed hoppers through a preheating zone (1450 mm long) prior to entering the reactor. The temperature inside the preheating zone was controlled by two kW vertical tube furnaces (each 725 mm long x 70 mm The flowrate of solids was controlled by means of a solids metering valve linked to a digital display and controlling device. The flowrate was measured at the entrance and exit of the reactor. After passing through the reactor, the solids were cooled before being collected in the solids receivers.

25 Mass flow controllers were used to regulate the flow of feed gas (methane). Gas passed in and out of the reactor via gas distributors and flowed through the reactor co-currently with the flow of solids. The reaction zone was 500 mm long. Product gas was sampled and analysed by an on-line gas chromatography (GC) unit for the duration of the experiment. There was also an option to continuously collect the product gas in a gas bag. Water traps were used to trap water originally present as moisture in the: cnbusted shale a 4 t n ;0 ,"i II 0d l~ c a C gn\02\12738.c 4 3 93 r i i 11 and water produced by the reaction of iron oxides with the product hydrogen.

The experimental conditions and results are shown below: Run Temperature Catalyst Catalyst Methane number 0 C) weight time residence conversion (kg.h/kmol) time (min) 1 900 150 5.0 34.1 2 900 75 5.8 19.3 3 940 75 5.4 31.5 Figure 3 shows a simplified flow sheet of the conventional oil shak, treatment plant. In Figure 3, the unit operations relating to the production of hydrogen by conventional steam reforming of the light hydrocarbons have been shown. As can be seen from Figure 3, dry shale enters retort 100 where the shale is heated. The retort off-vapour is sent to oil recovery section 110 to recover raw oil. The retort gas leaving oil recovery section 110 is treated to remove ammonia, hydrogen sulphide, carbon dioxide and C5+ hydrocarbons (120). Natural gas and hydrotreater gas are added to the retort gas and the gas mixture, called the hydrogen plant feed gas (122), is used as feed stock for the 6 hydrogen plant. This gas mixture comprising hydrogen, carbon monoxide and

C

1

C

5 hydrocarbons is fed into steam reformer 130 where conversion to hydrogen and carbon monoxide occurs. Further production of hydrogen takes place in water gas shift converter 132. Carbon dioxide is removed in vessel 134 and any remaining carbon oxides are converted to methane in methanator 136.

0$ '4 The gas stream comprising principally hydrogen and methane is used to treat the raw oil in hydrotreater 140 to produce a synthetic crude oil.

Spent shale from retort 100 is fed to combustor 115 where it is combusted to produce heat and combusted shale. The combusted shale is partly recycled to the retort, with the remainder being sent for sensible heat recovery prior to disposal.

Figure 4 shows a modified process for the treatment of oil shale in which the hydrogen required for hydrotreating of the raw oil is produced according to gn\02\12738.c 43 93

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ii -12the invention. In the process shown in Figure 4, dry shale is fed to retort 200 and raw oil is removed from the retort off-vapour in oil recovery section 210.

As in the conventional process, the retort gas from the oil recovery section 210 is cleaned in gas treating unit 220 and mixed with natural gas and hydrotreater gas to form a gas mixture, called hydrogen plant feed gas (222).

Spent shale from the retort is fed to combustor 215 and combusted to produce heat and combusted shale. Air 215, which may be oxygen enriched or oxygen reduced, is also fed to the combustor. The hot combusted shale is contacted with the hydrogen plant feed gas in catalytic dehydrogenator 230 to produce hydrogen. Carbon produced by the dehydrogenation reaction is laid down upon the combusted shale. A small shift converter 232 produces further hydrogen from carbon monoxide produced from the retort and the hydrogen rich gas stream, called make-up hydrogen stream (234), is then fed to the hydrotreater 240 to treat the raw oil and produce synthetic crude oil.

Combusted shale from catalytic dehydrogenator 230 is partly recycled to retort 200, with the remainder being sent for heat recovery and disposal. Catalytic dehydrogenator 230 may also include gas recycle streams and gas separation plant (not shown) to ensure that the product gas has sufficient hydrogen purity Hto meet the requirements of hydrotreate. feed. These extra items of plant may 20 not be required when high conversion is achieved in the dehydrogenator.

As can be seen by comparing Figure 4 with Figure 3, production of 4 hydrogen by contacting hot combusted shale with a gaseous mixture of treated retort gas, natural gas and hydrotreated gas simplifies the process. The °combusted shale has three roles: heat carrier, dehydrogenation catalyst and carbon collector. In the modified process of Figure 4, a compact dehydrogenator and a small shift converter replace a large steam reformer and shift converter, a carbon dioxide scrubber and a methanator.

Preliminary mass/energy balances for the standard and modified processes (50,000 barrels per day) were performed for comparison. Table 1 gives a summary of specified operating conditions of the two processes. Table 2 gn\02\12738.c 4 3 93 13 oo ooo mao ai Do n or o o, o o oo o compares their features. ~tailed results are given in Tables 3 and 4.

The only differences in the operating conditions of the two processes are retort and combustor temperatures. Although both processes recycle shale at the same temperature, a higher carbon content and hence a higher sensible heat of recycled shale from the modified process of the present invention leads to a higher retort temperature.

Tables 1 and 4 show that a modest drop iii temperature of recycled shale (37°C) is adequate to provide process heat to the shale dehydrogenator. Indeed, the heat load of the dehydrogenator is very low, only about one-fifth that of the reformer (Table 2).

The major economic benefit of the modified process is replacement of the very expensive H 2 plant with a compact shale dehydrogenator and a small shift converter (one-third of the original capacity) which is used to remove carbon monoxide produced in the retorting step.

Due to the difference in the form of rejected carbon -CO 2 (steam reforming) vs. carbon deposited on recycled shale (dehydrogenation) the modified process of the example emits 20% less total carbon dioxide, which is a green house gas. Higher content (4.6 vs. of finely dispersed carbon in the recycled shale could result in a reduction in oil coking reactivity and hence 20 higher oil yield. Exit shale which also has a high content of finely dispersed carbon could be very suitable for co-disposal with retort water. Retort gas contains up to 0.5 vol% SO 2 Calcium oxide in combusted shale could remove this environmental concerning gas effectively and economically.

The main drawback of the modified process is a higher requirement for feed stock gas. However, the cost of imported natural gas is low and would easily be offset by the anticipated significant reduction in operating costs associated with deleted operating units. Although the combustor heat load is about 6% higher, higher carbon content shale leads to a lower required carbon combustion conversion. The combustor size therefore would probably be unchanged. Hydrogen purity of the modified process depends on conversion of gn\02\1 273 8 .c 4 3 93

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i 14hydrocarbon gases. An acceptable purity of 95% can be obtained at a conversion of World hydrogen production for ammonia synthesis in 1988 was about million t/y (350 kg/s) which accounts for more than 50% world total hydrogen production. Hydrogen production for petroleum refining accounts for about An ammonia synthesis plant integrated with a shale-derived syncrude plant could result some capital cost savings by sharing a common shale dehydrogenating stage. Cheap hydrogen for refining syncrude could also be obtained from an integrated plant to produce transport fuels from oil shale.

1, r rr a TABLE 1 OPERATING CONDITIONS PROCESS STANDARD INVENTION Wet shale moisture, dry basis 25 Shale temperature, °C 300 300 Recycled shale temperature, °C 850 850 Recycle ratio 1.5 Retort temperature, oC 529 538 Combustor temperature, °C 850 887 Exit flue gas temperature, °C 200 200 Exit shale temperature, °C 250 250 Syncrude H/C 2.0 gn\02\12738.c 4 3 93 15 TABLE 2

COMPARISON

PROCESS STANDARD INVENTION Capacity of operation units, kg/s Steam reformer (H 2 targeted) 2.88 0 Shale dehydrogenator (H 2 targeted) 0 2.88 Shift converter (CO) 6.4 2.1

CO

2 removal (CO 2 18.2 0 Methanator (total gas) 7.5 0 Heat load (reformer/dehydrogenator), MW 182 36 Total CO 2 emission, kg/s 131 107 Combustor heat load, MW 828 875 Carbon combustion, 76 62 Carbon in recycled/exit shale, 2.3 4.6

H

2 purity, vol% 98.5 95.0 Retort/hydrotreater gas used, kg/s 17.0 19.3 Natural gas import, kg/s 0 oo a iaa o oc :oo o ecoe onn~

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16- O 0 .0 000 0 TABLE3 MASS AND ENERGY BALANCES FOR STANDARD RECYCLE SHALE PROCESS SOLIDS NAME: Wet Dry Spent Combusted Recycl Exch Exit Shale Shale Shale Shale e Shale Shale Shale Temperature, *C 25.0 300.0 528.5 850.0 850.0 850.0 250.0 Enthalpy, MW 0.0 214.1 683.5 1109.6 758.0 351.6 104.4 Total Flow, Kg/s 628.25 502.60 378.35 1103.55 753.90 349.65 349.65 Ash 341.87 341.87 341.87 1079.00 737.13 341.87 341.87 Korogen 137.21 137.21 0.00 0.00 0.00 0.00 0.00 Char 0.00 0.00 31.78 24.55 16.77 7.78 7.78 Moisture 125.65 0.00 0.00 0.00 0.00 0.00 0.00 Bound-H 2 0 7.54 7.54 1.51 0.00 0.00 0.00 0.00 Bound-CO 2 15.98 15.98 3.20 0.00 0.00 0.00 0.00 GAS Retort NAME: Recycle Combust'n Exit Off- Air Fluegas FBC Air Gas Fluegas Fluegas Vapour Temperature, 0 C 25.0 200.0 446.0 850.0 200.0 200.0 528.5 Enthalpy, MW 0.0 310.3 557.9 959.5 746.0 435.7 212.8 Total Flow, kg/s 313.13 330.55 643.68 668.93 794.58 464.04 124.25 1-12 0.00 0.00 0.00 0.00 0.00 0.00 0.29 Steam 0.00 96.73 96.73 106.88 232.53 135.79 20.55

H

2 S 0.00 0.00 0.00 0.00 0.00 0.00 0.17 CO 0.00 0.00 0.00 0.00 0.00 0.00 2.14

CO

2 0.00 62.40 62.40 149.99 149.99 87.59 11.07 Methane 0.00 0.00 0.00 0.00 0.00 0.00 2.04 Ethylene 0.00 171.42 0.00 0.00 0.00 0.00 2.14 Ethane 0.00 0.00 0.00 0.00 0.00 0.00 1.26 Propylene 0.00 0.00 0.00 0.00 0.00 0.00 1.77 Propane 0.00 0.00 0.00 0.00 0.00 0.00 0.71 1-Butene 0.00 0.00 0.00 0.00 0.00 0.00 1.03 n-Butane 0.00 0.00 0.00 0.00 0.00 0.00 0.36 Oil vapour 0.00 0.00 0.00 0.00 0.00 0.00 80.75 02 72.49 0.00 72.49 0.00 0.00 0.00 0.00

N

2 240.65 0.00 412.07 412.07 412.07 240.65 0.00 GAS NAME: Retort Hydrotr H2plant H2plant Shift Makeup Hydrotr Gas Gas Feed Gas Out In H2 H2 In Temperature *C 25.0 25.0 NA NA 350.0 25.0 25.0 Enthalpy, MW 0.0 0.0 NA NA 12.5 0.0 0.0 Total Flow, kg/s 23.98 4.39 8.06 38.34 38.34 3.36 8.38

H

2 0.29 0.26 0.30 2.89 2.89 2.88 5.90 Steam 1.03 1.95 0.61 16.70 16.70 0.00 0.10

H

2 S 0.17 0.43 0.00 0.00 0.00 0.00 0.02 CO 2.14 0.00 1.16 0.13 0.13 0.00 0.00

CO

2 11.07 0.00 0.00 18.15 18.15 0.00 0.00 Methane 2.04 0.16 1.19 0.10 0.10 0.10 0.39 Ethylene 2.14 0.10 1.21 0.10 0.10 0.10 0.22 Ethane 1.26 0.63 1.03 0.08 0.08 0.08 0.86 Propylene 1.77 0.08 1.00 0.08 0.08 0.08 0.15 Propane 0.71 0.73 0.78 0.06 0.06 0.06 0.66 1-Butene 1.03 0.05 0.58 0.05 0.05 0.05 0.07 n-Butane 0.36 0.02 0.20 0.02 0.02 0.02 0.02 gn\02\127 3 8 .c 3 3 93 17 TABLE 4 MASS AND ENERGY BALANCES FOR MODIFIED RECYCLE SHALE PROCESS OF TIE PRESENT INVENTION SOLIDS NAME: Wet Dry Spent Combusted Recycl Exch Exit Shale Shale Shale Shale e Shale Shale Shale Temperature, 'C 25.0 300.0 538.5 887.0 850.0 850.0 250.0 Enthalpy, MW 0.0 214.1 709.4 1196.1 787.1 373.5 107.8 Tota' Flow, Kg/s 628.25 502.60 378.35 1102.19 753.90 357.73 357.73 Ash 341.87 341.87 341.87 1062.34 720.48 341.87 341.87 Korogen 137.21 137.21 0.00 0.00 0.00 0.00 0.00 Char 0.00 0.00 31.78 39.84 33.42 15.86 15.86 Moisture 125.65 0.00 0.00 0.00 0.00 0.00 0.00 Bound-H 2 0 7.54 7.54 1.51 0.00 0.00 0.00 0.00 Bound-CO 2 15.98 15.98 3.20 0.00 0.00 0.00 0.00 GAS Retort NAME: Recycle Combust'n Exit Off- Air Fluegas FBC Air Gas Fluegas Fluegas vapour Temperature, 'C 25.0 200.0 481.5 887.0 200.0 200.0 538.0 Enthalpy, MW 0.0 257.5 522.9 911.5 697.8 440.3 216.2 Total Flow, kg/s 330.88 282.42 613.30 639.71 765.36 482.94 124.25 H2 0.00 0.00 0.00 0.00 0.00 0.00 0.29 Steam 0.00 79.70 79.70 90.33 215.98 136.28 20.55 H1 2 S 0.00 0.00 0.00 0.00 0.00 0.00 0.17 CO 0.00 0.00 0.00 0.00 0.00 0.00 2.14

CO

2 0.00 54.02 54.02 146.40 146.40 92.38 11.07 Methane 0.00 0.00 0.00 0.00 0.00 0.00 2.04 Ethylene 0.00 148.70 0.00 0.00 0.00 0.00 2.14 Ethane 0.00 0.00 0.00 0.00 0.00 0.00 1.26 Propylene 0.00 0.00 0.00 0.00 0.00 0.00 1.77 Propene 0.00 0.00 0.00 0.00 0.00 0.00 0.71 1-Butene 0.00 0.00 0.00 0.00 0.00 0.00 1.03 n-Butane 0.00 0.60 0.00 0.00 0.00 0.00 0.36 Oil vapour 0.00 0.00 0.00 0.00 0.00 0.00 80.75 02 76.59 0.00 76.59 0.00 0.00 0.00 0.00

N

2 254.28 0.00 402.99 402.99 402.99 254.28 0.00 GAS NAME: Retort H2drotr H2plant H2plant Shift Makeup Hydrotr Gas Gas Feed Gas Out In H2 H2 In Temperature 'C 25.0 25.0 558.0 870.0 350.0 25.0 25.0 Enthalpy, MW 0.0 0.0 29.9 46.4 16.6 0.0 0.0 Total Flow, kg/s 23.98 7.79 19.31 9.86 9.86 6.76 15.88 H2 0.29 0.26 0.55 2.73 2.73 2.88 5.91 Steam 1.03 1.95 1.12 1.12 1.12 0.00 0.10

H

2 S 0.17 0.43 0.00 0.00 0.00 0.00 0.02 CO 2.14 0.00 2.14 2.14 2.14 0.00 0.00

CO

2 11.07 0.00 0.00 0.00 0.00 0.00 0.00 Methane 2.04 1.11 4.20 1.05 1.05 1.05 3.11 Ethylene 2.14 0.71 2.85 0.71 0.71 0.71 1.58 Ethane 1.26 1.15 2.41 0.60 0.60 0.60 2.01 Propylene 1.77 0.59 2.36 0.59 0.59 0.59 1.07 Propane 0.71 1.12 1.82 0.46 0.46 0.46 1.37 1-Butene 1.03 0.34 1.37 0.34 0.34 0.34 0.51 n-Butane 0.36 0.12 0.47 0.12 0.12 0.12 0.18 3 3 93 gn\02\12738.c

Claims (24)

1. A process for the production of hydrogen by dehydrogenation of light hydrocarbons comprising contacting the light hydrocarbons with combusted oil shale under conditions suitable for a dehydrogenation reaction to occur.

2. A process as claimed in claim 1 wherein the contacting of the light hydrocarbons with the combusted oil shale takes place at a temperature of from 500°C to 1200 0 C.

3. A process as claimed in claim 2 wherein the temperature is from 750°C to 950 0 C.

4. A process as claimed in any one of claims 1 to 3 wherein the contacting of the light hydrocarbons with the combusted oil shale takes place at a pressure of from 0.1 to 50 atmospheres.

5. A process as claimed in claim 4 wherein the pressure is from 1 to 20 atmospheres.

6. A process as claimed in any one of claims 1 to 5 wherein the process is operated at a catalyst weight time of from 1 20000 kg of catalyst.hour per kmol of light hydrocarbons.

7. A process as claimed in claim 6 wherein the catalyst weight time is from 10 2000 kg of catalyst.hour per kmol of light hydrocarbons.

8. A process as claimed in any one of the preceding claims wherein the process is operated at a combusted shale residence time of from 0.5 to 120 minutes.

9. A process as claimed in claim 8 wherein the combusted shale residence time is from 3 to 30 minutes. gn\02\12738.c 33 93 ft,-- p.- 19 A process as claimed in any one of the preceding claims wherein the combusted shale has a carbon content of from 0 to 20% by weight.

11. A process as claimed in claim 10 wherein the combusted shale has a carbon content of from 1.5 to 8% by weight.

12. A process as claimed in any one of claims 1 to 11 wherein the light hydrocarbons is selected from the group comprising natural gas, retort gas, hydrotreater gas and mixtures thereof.

13. A process for the production of synthetic crude oil from oil shale comprising: supplying oil shale to a retort; ii) heating said shale in said retort to produce retort off-vapour and spent shale; (iii) combusting said spent shale in a combustor to produce heat and combusted shale; (iv) passing said retort off-vapour through an oil recovery stage to separate raw oil from retort gas; and upgrading said raw oil to synthetic crude oil by treating said raw oil with hydrogen in an hydrotreating process; wherein said hydrogen is produced by the dehydrogenation of a light hydrocarbons-containing gas in the presence of said combusted shale.

14. A process as claimed in claim 13 wherein said spent shale is combusted at a temperature of from 500 to 1500°C, preferably 750 1000°C. gn\02\12738.c 3 3 93 20 A process as claimed in claim 13 or claim 14 wherein said spent shale is combusted at a pressure ranging from 0.1 to 50 atmospheres, preferably 1 20 atmospheres.

16. A process as claimed in any one of claims 13 to 15 wherein an inlet gas fed to said combustor has a low oxygen content.

17. A process as claimed in claim 16 wherein the oxygen content of the inlet gas is from 2.0 to 21.0%.

18. A process as claimed in any one of claims 13 to 17 wherein the combusted shale has a residence time in the combustor of from 10 seconds to 2 hours.

19. A process as claimed in claim 18 wherein the residence time of the combusted shale in the combustor is from 30 to 600 seconds. A process as claimed in any one of claims 13 to 19 wherein said hydrogen is produced by contacting said light hydrocarbons with said combusted shale, said contacting taking place at a temperature of from 500 to 1200 0 C.

21. A process as claimed in claim 20 wherein said temperature is in the range of 750 to 950°C.

22. A process as claimed in any one of claims 13 to 21 wherein the contacting of the light hydrocarbons with the combusted shale takes place at a pressure of from 0.1 to 50 atmospheres.

23. A process as claimed in claim 22 wherein the pressure is from 1 to 20 atmospheres. 24 A process as claimed in any one of claims 13 to 23 wherein the dehydrogenation of the light hydrocarbons-containing gas takes place at a gn\02\12738.c 3 3 93 MM-M-000 F- 21 catalyst weight time of from 1 20000 kg of catalyst.hour per kmol of light hydrocarbons, preferably 10 2000 kg of catalyst.hour per kmol of light hydrocarbons. A process as claimed in any one of claims 13 to 24 wherein the hydrogenation of the light hydrocarbons-containing gas takes place with a catalyst residence time of from 0.5 to 120 minutes, preferably 3 to 30 minutes.

26. A process as claimed in any one of claims 13 to 25 wherein said light hydrocarbons-containing gas is selected from the group comprising retort gas, hydrotreater gas, natural gas and mixtures thereof.

27. A catalyst for the catalytic dehydrogenation of light hydrocarbons comprising combusted oil shale.

28. Synthetic crude oil produced by the method as claimed in any one of claims 13 to 26.

29. A process for producing hydrogen substantially as hereinbefore described with reference to Examples 1 to 3. A process for the production of synthetic crude oil from oil shale substantially as hereinbefore described with reference to Figure 4. DATED March 5, 1993 CARTER SMITH BEADLE Fellows Institute of Patent Attorneys of Australia Patent Attorneys for the Applicant: COMMONWEALTH SCIENTIFIC AND INDUSTRIAL RESEARCH ORGANISATION gn\02\12738.c 5 3 93 i; i I 22 ABSTRACT A process for the production of hydrogen by the dehydrogenation of light hydrocarbons comprises contacting the light hydrocarbons with combusted oil shale. The reaction may take place at a temperature of 500 1200°C. The combusted oil shale is effective to catalyse the dehydrogenation reaction. The process may be incorporated into an oil recovery process for producing synthetic crude oil from oil shale. The hydrogen produced by the process is used in a hydrotreating operation in which raw oil recovered from oil shale is upgraded to synthetic crude oil. r s gn\02\12738.c 5393 i.-