Producing liquid hydrocarbon streams by hydrogenation of fossil-based feedstock

An economic route to transport fuels from low grade feedstocks containing organic polycyclic components and mineral and/or metallic components comprises separating the feedstock into (a) a residue containing fuel values and substantially all of the mineral and metallic components and (b) a liquid hydrocarbon fraction; hydrogenating the liquid hydrocarbon fraction; providing hydrogen for the hydrogenation by steam reforming methane-containing gas recovered from the hydrogenated material; and providing the heat for the steam reforming by immersing the reformer reactor tubes in a fluidized bed heated by combustion of residue from the separation step. If the fluidized bed is pressurized, the process can be made substantially self-contained with all the heat and power requirements for steam generation and feedstock compression for hydrogenation and steam reforming being recovered from the flue gas.

This invention relates to the treatment of a fossil-based feedstock 
containing a substantial proportion of high molecular weight organic 
polycyclic components to recover a more valuable liquid hydrocarbon stream 
e.g. suitable for use as or in the production of transport fuels such as 
gasoline, diesel oil, gas oils, kerosine, aviation gasoline and aviation 
turbine fuel. 
High molecular weight organic polycyclic components are found in 
substantial quantities in crude oils, in which context asphaltenes form a 
significant proportion of these polycyclic components, and are 
concentrated in the residues of such oils after distillation. They are 
also present in substantial amounts in solid fossil fuels such as coal, 
peat, lignite and shale, and also in tar sands. 
Such polycyclic components are often heterocyclic and can include in their 
structures atoms of oxygen, sulphur and/or metals. 
Examples of fossil-based feedstocks with which this invention is concerned 
are, thus, oil-based feedstocks such as crude oils, distillation residues 
of crude oils such as atmospheric residues and vacuum residues, 
vis-breaking residues, cracking residues and oils derived from the 
distillation of shale, and solid materials such as coal, lignite, peat, 
shale and tar sands. 
It is known that valuable liquid hydrocarbon streams such as are suitable 
for use as or in the production of transport fuels may be derived from 
such fossil-based feedstocks by hydrogenation. 
In one method, the hydrogen for the hydrogenation is generated by gasifying 
a portion of the fossil-based feedstock or a residue derived from it such 
as the char formed in the solvent extraction of a coal. However, the 
capital costs of gasification plants are so high that even with the 
benefit of using a cheap feedstock such as char for the hydrogen, the 
process has rarely been commercially employed. 
In another method, such as described in BP Nos. 1289158 and 1525436, it is 
proposed to produce hydrogen for the hydrogenation by steam reforming a 
gaseous hydrocarbon stream which is formed as a byproduct of the 
hydrogenation. The capital costs of steam reforming can be less than half 
those of gasification but premium gaseous or high grade liquid fuels are 
required to heat the reformer. This is because with the known heater 
designs, the accurate control of reformer tube temperatures, which is 
essential for the avoidance of premature tube failure and/or carbon 
build-up on the catalyst, can only be achieved if such fuels are employed, 
and even then optimum reaction conditions cannot be ensured. 
In practice this means that either a part of the gaseous hydrocarbon stream 
must be employed as the fuel or a separate source of supply of the fuel is 
required. In the former case, the balance of the gaseous hydrocarbon 
stream is insufficient to produce the required amount of hydrogen and in 
the latter case, the capital cost advantage of steam reforming over 
gasification stands to be offset largely or wholly by the higher cost of 
the fuel. 
Whichever alternative is adopted, therefore, producing the hydrogen 
accounts for a large proportion of the overall cost of the process and can 
account for as much as 30% to 50% of the value of the liquid hydrocarbon 
product. For this reason, the industry has recognised for many years the 
importance of this step of the process and much effort has been directed 
at reducing the expense of it. For example, attention has been directed to 
improving the selective utilisation of the hydrogen to the production of 
the desired products, to reducing the capital cost of the gasification 
route, to improving heat recovery from the steam reforming step, and even 
to re-siting refinery operations. However, although the steam reforming 
route still has a major capital cost advantage over the known 
alternatives, no attention appears to have been directed to reducing costs 
by use of a low grade fuel to provide the heat for the reforming and no 
practicable proposal has been put forward. 
According to the present invention there is provided a process which 
comprises the steps of separating a fossil-based feedstock containing a 
substantial proportion of high molecular weight organic polycyclic 
components and also mineral and/or metallic components to provide (a) a 
residue containing fuel values and substantially all of said mineral 
and/or metallic components and (b) a liquid hydrocarbon fraction; reducing 
the average molecular weight of said liquid hydrocarbon fraction by 
hydrogenation; fractionating hydrogenation material thereby obtained to 
form a gaseous fraction containing substantially all the methane in said 
material and at least one liquid hydrocarbon stream; and providing 
hydrogen for said hydrogenation by steam reforming a methane-containing 
gas provided from said gaseous fraction, the reforming being effected at 
elevated temperature in a reactor vessel which is at least partly immersed 
in a fluidised bed of finely divided solid material which is heated by 
combustion of a fuel provided at least in part, and preferably entirely or 
substantially entirely, from said residue of said separation step. 
By means of the invention, the cost of providing the heat for the stream 
reforming step, which cost forms a very significant proportion of the cost 
of producing the hydrogen, is substantially reduced without very much 
change in the capital cost of the steam reforming equipment, thereby 
providing an important and substantial saving in the overall cost of the 
production of the liquid hydrocarbons. 
Other advantages also accrue from the use of a fluidised bed combustor; in 
particular, improved temperature control through more uniform transfer of 
heat to the reformer tube walls, elimination of the need to ensure an 
exact distribution of feed to each reformer tube, higher heat transfer 
rates with consequent reduction in the temperature of the flue gas leaving 
the reformer tube zone, lower temperature differences between heat source 
and the tube wall, greater freedom in the shape and orientation of the 
reactor tubes, more compact arrangement of tubes and consequential 
reduction in overall apparatus size and refractory requirements, lower 
combustion flame temperature with consequential reduction in concentration 
of oxides of nitrogen in the flue gas, possibility of upward or downward 
flow of reaction mixture through the reactor tubes and, with the former, 
of using fluidised catalyst beds in the tubes, and superior turn down and 
flexibility of process control. 
A particular advantage of the process of the invention, however, is that 
the fluidised bed may be under superatmospheric pressure and thus the 
pressure drop across the reformer reactor tube walls may be reduced or 
eliminated thereby permitting extended tube life, higher operating 
temperatures, use of thinner tube walls, use of lower grade tube material, 
higher pressures in the reformer tubes or a combination of two or more of 
these possibilities. It also provides the possibility of using ceramic 
tubes under compression, in which mode they offer better performance. 
A further advantage of operating the fluidized bed combustor at 
superatmospheric pressure is that power for the process of the invention 
may be provided by expanding the flue gas from the combustor through an 
expansion engine such as a turbine. The extraction of energy from the flue 
gas in this manner is more efficient than the conventional raising of high 
pressure superheated steam and the subsequent use of the steam to drive 
compressors or electricity generators. Moreover, by suitable choice of the 
amount of excess oxidant gas supplied to the combustor and of the 
presssure of the combustor, all or substantially all the power 
requirements of the process, and in particular for compressing the oxidant 
gas for the fluidised bed combustor, and compressing gases and pumping 
liquids in the process, e.g. for the hydrogenation and steam reforming 
steps, may be satisfied in this manner. 
Thus, the process may be substantially self-contained, with all the fuel 
requirements for the steam reformer being provided by the residue from the 
separation step, all the hydrogen for the hydrogenation being provided 
from the gaseous material produced by the hydrogenation and all the energy 
requirements being provided by expansion of the flue gas from the 
fluidised bed combustor. 
The feedstock for the process may be a crude oil, a liquid residue derived 
from crude oil, an oil derived from tar sands, or a solid, which may be an 
oil residue or a solid fossil fuel such as coal or an oil-bearing shale or 
a tar sand. The nature of the separation step will be determined by the 
nature of the feedstock. The following Table indicates possible separation 
steps for various feedstocks which may be used in the process of this 
invention. However, it is to be understood that the feedstocks which are 
suitable for use in the process of this invention are not limited to those 
listed in the Table and the processes indicated as suitable for use for 
separating these feedstocks in accordance with the first step of the 
process of the invention do not cover all the possible processes. Other 
feedstocks and separation processes will not be apparent to those skilled 
in the art. 
So far as concerns oil-based materials, the invention is particularly 
suited to the treatment of heavy crude oils, producable tar oils such as 
Orinoco Tar which has API gravities in the order of 8 to 12, and crude oil 
residues having API gravities in the same range. Coal, lignite and tar 
sands are also particularly suitable feedstocks. 
TABLE 
______________________________________ 
Feedstock Separation Step 
______________________________________ 
crude oil (a) 1 or 2 or 1+3 
shale oil (b) (1 or 2 or 1+3) + (4 or 5 or 6 or 7) 
oil from tar (c) (1 or 2 or 1+3) + (4 or 5) + (6 or 7) 
sands 
residue of atmospheric 
(a) 3 or 4 or 5 or 6 or 7 
distillation of crude oil 
(b) 3 + (4 or 5 or 6 or 7) 
(c) 3 + (4 or 5) + (6 or 7) 
(d) (4 or 5) + (6 or 7) 
residue of vacuum 
(a) 4 or 5 or 6 or 7 
distillation of crude oil 
(b) (4 or 5) + (6 or 7) 
vis-breaker residue 
6 or 7 
cracking residue 
coal (a) solvent extraction (which term includes 
lignite supercritical solvent extraction and 
peat solvent extraction with simultaneous 
hydrogenation e.g. by hydrogenating a 
coal-in-oil slurry with recycle of 
hydrogenated oils) 
(b) carbonisation (including flash 
carbonisation) 
oil-bearing shale distillation or solvent extraction 
tar sands separation of oil or distillate by paraf- 
finic solvent extraction or by coking or 
by distillation, to produce a fuel- 
containing sand residue and a liquid 
hydrocarbon fraction 
______________________________________ 
KEY 
1. Atmospheric distillation 
2. Vacuum distillation 
3. Vacuum distillation of the residue from atmospheric distillation 
4. Vis-breaking the residue of atmospheric or vacuum distillation 
5. Thermal cracking the residue of atmospheric or vacuum distillation 
6. Coking the residue of atmospheric or vacuum distillation or the residue 
or vis-breaking or of cracking 
7. Deasphalting the residue of atmospheric or vacuum distillation or the 
residue of vis-breaking or of cracking 
The terms "atmospheric distillation" and "vacuum distillation" include both 
single- and multi-stage distillations. 
The separation may also comprise hydrocracking oil with an ebullated bed 
operated in such fashion as to leave asphaltene substantially unreacted, 
e.g. as in the H-Oil process. 
The immediate liquid product of the separation step may be subjected to 
further treatment, e.g. fractionation, to provide the liquid hydrocarbon 
fraction to be hydrogenated. For example, where the feedstock for the 
process is a solid substantially infusible fuel such as coal, lignite or 
peat and the separation step includes solvent extraction, it will 
generally be desirable to distil the liquid product to recover solvent for 
recycle. 
In general, no further processing of the distillate from the step of 
separating oil-based feedstocks will be required but a further processing 
step is not excluded. 
Thus, to summarise, the step of separating the feedstock into the residue 
and the liquid hydrocarbon fraction may comprise a single step, e.g. 
atmospheric or vacuum distillation, or a series of process steps with the 
residue and the liquid hydrocarbon fraction being the products of 
different steps of the operation; e.g. as in the distillation of crude oil 
followed by deasphalting the distillation residue, wherein the liquid 
hydrocarbon fraction may be provided from the distillation (optionally 
after further fractionation) and the residue may be provided from the 
deasphalter, or as in the solvent extraction of coal, peat or lignite, 
where the residue may be the solid residuum of the solvent extraction and 
the liquid hydrocarbon fraction is provided from the solution after 
treatment to recover the solvent. 
All the processes employed singly or in combination to effect the 
separation in accordance with the invention may be conducted in accordance 
with well known principles. The nature of the residue and of the liquid 
hydrocarbon fraction can vary widely and will depend upon the nature of 
the feedstock and of the process or processes employed to effect the 
separation. For example, where the feedstock is crude oil or is derived 
from crude oil, the residue from atmospheric or vacuum distillation or 
from cracking or vis-breaking may be a more or less viscous liquid whereas 
the residue from coking or deasphalting will normally be a tarry or 
asphaltic solid or semi-solid at ambient temperatures. The residue from 
the carbonisation or solvent extraction of coal, lignite or peat will be a 
carbonaceous solid. 
The liquid hydrocarbon fraction employed for the hydrogenation may vary 
widely in composition, viscosity and average molecular weight, depending 
on the nature of the feedstock and kind of process or processes employed 
for the separation. More than one liquid hydrocarbon fraction may be 
produced by the separation and one or more of the fractions may be 
hydrogenated. Where two or more fractions are to be hydrogenated, the 
hydrogenations may be effected in separate hydrogenation reactors. For 
example, where the separation includes distillation of oil employing 
atmospheric and vacuum distillations, fractions from both these 
distillations may be separately hydrogenated. 
The second step of the process of the invention comprises hydrogenation of 
the liquid hydrocarbon fraction to reduce its average molecular weight. 
The hydrogenation may be conducted in known manner and employing 
well-established conditions. In general it involves contacting the liquid 
hydrocarbon fraction with hydrogen at elevated temperature and 
superatmospheric pressure, optionally in the presence of a catalyst. For 
example, the hydrogenation may comprise hydrotreatment, hydrocracking or 
hydrodesulfurisation. Hydrocracking tends to have the highest demand for 
hydrogen but is often the preferred procedure. 
The hydrogenation may involve more than one hydrogenation step, for example 
the hydrogenation may comprise a first hydrogenation step followed by 
fractionation of the first hydrogenation product and further hydrogenation 
of one or more of the fractions so produced. Further processing steps such 
as e.g. cracking may be interspersed between hydrogenation steps. 
In the next step, hydrogenated material recovered from the hydrogenation, 
and which may comprise a part or all of the product of the hydrogenation 
step, is fractionated to produce a gaseous fraction containing methane and 
at least one liquid hydrocarbon stream. The fractionation may be carried 
out in known manner and employing well-known conditions. The gas may 
include C.sub.2, C.sub.3 and possibly C.sub.4 hydrocarbons but it may be 
preferred to recover at least the C.sub.3 and C.sub.4 components as a 
separate cut e.g. for use as LPG and bottled gas. The gas may also include 
unreacted hydrogen and additionally hydrogen sulfide if the liquid 
hydrocarbon fraction which is hydrogenated contains sulfurous materials. 
The fractionation may be operated in one or more steps and to produce a 
single liquid hydrocarbon product or a plurality of liquid cuts, e.g. 
suitable for use as or in the production of motor spirits (gasoline), 
aviation turbine fuel, aviation gasoline, vaporising oil, kerosine, diesel 
oil, heavy diesel oil, etc. 
Where the liquid hydrocarbon fraction to be hydrogenated has been produced 
by distillation, e.g. as in the atmospheric and/or vacuum distillation of 
oil, the fractionation of the hydrogenation product may be effected by 
recycling the product to said distillation. Alternatively, it may be 
fractionated in a separate step. Where it is recycled, the recovery of the 
methane-containing gas will generally be effected by a degassing step 
prior to recycling. 
Where there is more than one hydrogenation step, it will generally be 
desirable to recover methane-containing gas from the product of each 
hydrogenation. 
In accordance with the invention, at least a part and preferably all of the 
hydrogen requirement of the hydrogenation step is obtained by 
steam-reforming methane-containing gas recovered from the hydrogenated 
material. The entire gaseous product of the hydrogenation may be steam 
reformed, if desired, but it may be treated first e.g. to separate 
specific fractions which may be required as premium products, as mentioned 
above. 
If insufficient gas is available from the hydrogenation to satisfy the 
hydrogen requirements of the hydrogenation, the deficiency may be resolved 
by additionally steam reforming at least a portion of the lighter liquid 
hydrocarbons produced by hydrogenation, e.g. hydrocarbons having 4 to 7 
carbon atoms or a suitable naphtha fraction. 
It is preferred to remove any sulfur-containing gases prior to steam 
reforming. It may also be desired to remove any unreacted hydrogen; 
however, the increased heat burden on the steam reformer from retaining 
some or all of the hydrogen is tolerable because low grade fuel is used. 
The feed to the reformer may also include methane-containing gas produced 
as a direct product of the separation step, e.g. gas from a distillation, 
pyrolysis or cracking step employed in the separation. The steam reforming 
operation may be effected in known manner and employing well-known 
conditions. 
To maximise hydrogen production and to eliminate carbon monoxide which 
interferes with hydrogenation catalysts it is preferred to encourage the 
CO shift reaction in known manner, e.g. by spraying water into the gas 
recovered from the reformer, and the carbon dioxide may then be removed 
from scrubbing, also in known manner. 
Any carbon monoxide still remaining in the reformate may be eliminated by 
methanation and steam in the reformate may be condensed out. 
Hydrogen for the hydrogenation step may then be provided from the remaining 
gas which will comprise primarily hydrogen and some unreacted methane. 
As indicated above, the fractionation of the hydrogenated material may 
include recycling it to a distillation train from which is derived a 
gaseous fraction and liquid hydrocarbon fraction to be hydrogenated. By 
such recycle, a wide degree of control of the nature of the desired net 
liquid product from the fractionation may be achieved. 
In accordance with the invention, the heat for the steam reforming is 
provided by at least partly immersing the reformer reactor vessel (which 
term is to be understood to include a single reactor vessel or a plurality 
of such vessels, e.g. catalyst-packed reactor tubes) in a fluidised bed of 
finely divided solid material which is heated by combustion of a fuel 
provided from the residue produced by the separation step. The process can 
be adapted so that all the fuel requirements are provided by said residue, 
which may be solid, semisolid or liquid in nature. This may be achieved 
e.g. by control of the separation step to produce the required amount of 
residue. 
Where the residue is an ash-containing solid, it is provided in particulate 
form, preferably having a maximum dimension not exceeding 6 mm, and more 
preferably with an average particle size of 0.25 to 1.0 mm, and is 
normally supplied to the bottom of the bed, the ash being removed either 
continuously or discontinuously. The fuel may be fed in dry powder form or 
as a slurry e.g. in water. Where the fuel is fusible, e.g. as in the case 
of oil-derived residues, it may be first melted and then fed to the 
fluidised bed in molton form. 
Where the fuel is ash-less or the ash formed is insufficient to form the 
fluidised bed, inert material may be added to provide or contribute to the 
formation of the fluidised bed. Any suitable inert material may be used 
and examples are alumina and sand. If desired, dolomite or limestone may 
be included in the bed to reduce the level of oxides of sulfur in the flue 
gas where sulfur-containing fuels are employed. The inert material and/or 
limestone or dolomite should ideally have a maximum particle dimension of 
6 mm and an average particle size in the range 0.25 to 1.0 mm. 
Where fuels containing high metal contents and in particular heavy oil or 
oil residue fuels containing vanadium are used, it may be advantageous to 
introduce magnesium in the form of dolomite to suppress hot metal erosion 
due to vanadium, particularly in the region of the expander gas turbine. 
The bed is fluidised by gas, normally the oxidant gas, e.g. air, required 
for the combustion, and is generally supplied from below. 
It is desirable to operate the fluid bed combustor with excess air, 
preferably in the range 10 to 100% and most preferably in the range 20 to 
50% excess of stoichiometric. It is also possible to operate the combustor 
with a deficiency of air with provision for burning off combustible gases 
by the addition of air after the fluid bed combustor.

Referring to FIG. 1 of the drawings, reference numeral 2 is an atmospheric 
distillation unit, 4 is a vacuum distillation unit and 6 is a vacuum 
residue processing unit such as a solvent de-asphalter. 8, 10 and 12 are 
hydrogenation units for treating, respectively, atmospheric distillates, 
vacuum distillates and solvent refined product from vacuum residue 
processing unit 6. 14 is an atmospheric distillation unit for processing 
hydrogenated products from hydrogenation units 8, 10 and 12, and 16 is a 
vacuum distillation unit for treating the residue from the atmospheric 
distillation unit. 20 is a refinery gas purification plant, 22 is a gas 
separation unit and 24 is a hydrogen production plant incorporating a 
residue-fired fluid bed reformer and the arrangement of which is shown in 
more detail in FIG. 3. 
Processing unit 6 could also be a vis-breaker or a coker. Other units such 
as catalytic reformers and catalytic crackers (not shown) may be 
incorporated in the overall processing unit which as a result would 
involve additional process streams not shown in the diagram. It will also 
be understood that variations in the illustrated flow arrangements may be 
made. 
It will also be understood that the units referred to above incorporate 
columns, vessels, pumps, heaters, etc, as is well known in the art, and 
that subsidiary process and service connections are not shown. It is also 
to be understood that the pipelines shown may represent a single pipeline 
or a plurality of pipelines in parallel. 
The feed to the plant may comprise crude oil supplied through line 102 
and/or a heavy oil or oil residue supplied through line 110 and/or line 
118. Feed supplied through line 102 is separated in atmospheric 
distillation unit 2 into a gaseous fraction removed overhead through line 
108, atmospheric distillate which is removed through line 104 and a 
residue which is removed through line 106. Provision is made for supply of 
some or all of this residue, if desired, as fuel for the hydrogen plant 24 
via valve 107 and pipelines 109 and 124. 
The remainder, if any, of the residue in line 106 and/or fresh feed 
supplied through line 110 is transferred via line 112 to vacuum 
distillation unit 4 where it is separated into a vacuum distillate removed 
in line 114 and a residue which is removed in line 116. There is provision 
for supplying some or all of this residue, if desired, as fuel for the 
hydrogen plant 24 via valve 117 and lines 119 and 124. 
Any remainder of the residue in line 116 and/or any fresh feed to the plant 
supplied through line 118 is transferred via line 120 to the vacuum 
residue processing unit 6 which, in the arrangement illustrated, is a 
solvent de-asphalting unit. Solvent de-asphalted product which has also 
been substantially stripped of mineral matter and its metal content such 
that it is a suitable general hydrogenation feed, is withdrawn from the 
unit via line 122 and the asphalt containing residue is withdrawn via line 
124. Some of the residue material in line 124 may be withdrawn via a valve 
or suitable control device 125 and line or transfer system 127. This 
residue may be used as a low grade fuel or asphalt or may be blended with 
light oil to produce a heavy fuel oil. The residue in line or alternate 
transfer system 124 is supplied as fuel to hydrogen plant 24. 
Atmospheric distillate in line 104 may be passed via line 132 for disposal 
as product or via line 130 to hydrogenation unit 8. Hydrogenated material 
from 8 may be disposed as product via line 133 and/or may pass via lines 
136 and 154 for distillation in unit 14. Surplus gas from unit 8 leaves 
via lines 134 and 152. 
Vacuum distillate in line 114 may pass via line 140 for disposal as product 
or via line 138 to hydrogenation unit 10. Hydrogenated material from 10 
may pass via line 143 for disposal as product and/or via lines 144 and 154 
for distillation in unit 14. Surplus gas is removed via line 142. 
De-asphalted oil in line 122 may pass via line 147 for disposal as product 
and/or via line 146 to hydrogenation unit 12. Hydrogenated material from 
12 will normally be passed via lines 150 and 154 for distillation in unit 
14 but some may be withdrawn as product (now shown). Surplus gas is 
recovered via line 148. 
The hydrogen requirements for hydrogenation units 8, 10 and 12 are provided 
at suitable pressure from line 184 via lines 210, 208 and 206, 
respectively. The source of the hydrogen is described below. 
The hydrogenated material collected in line 154 from units 8, 10 and 12 is 
passed to atmospheric distillation unit 14 where it is fractionated into 
one or more liquid fractions recovered in line 156, stripped gas which is 
recovered via line 158 and atmospheric residue which leaves via line 160 
and is passed to vacuum distillation unit 16. The liquid fraction or 
fractions from unit 16 are recovered in line 162 and may be disposed of as 
product via line 168 and/or recycled for further hydrogenation in hydrogen 
unit 10 via line 164. Residue from unit 16 is passed back to hydrogenation 
unit 12 via line 170. 
In this way, the distillates from units 2, 4 and 6 may be recirculated and 
upgraded by hydrogenation and subsequent distillation and further 
recirculation to produce enhanced lower molecular weight product 
ultimately recovered from 133, 143, 156 and/or 168. 
The gas from distillation units 2 and 14 and hydrogenation units 8, 10 and 
12 is collected in line 152 and passed to gas purification unit 20 in 
which undesirable constituents such as hydrogen sulphide and ammonia are 
removed via line 172. Purified gas is recovered in line 174 and passed to 
optional separation unit 22 which separates the gas into two or mre 
fractions. Specific fractions such as hydrogen and/or propane and/or 
butane may be removed via line 176, which may be a plurality of product 
pipelines, and may be disposed of as product or recycled to specified 
process e.g. the H.sub.2 into line 204. The remaining gas comprising the 
methane values in the gas stream and possibly other light hydrocarbons, is 
passed via line 178 to hydrogen plant 24. 
This plant comprises a steam reformer heated by a fluidised bed combustor, 
as more particularly illustrated in FIG. 3. The arrangement is described 
in more detail below. Fuel for this plant enters via pipeline or conveying 
system 124 and adsorbent and/or inert material to form the bed is provided 
by a transport system 186. Combustion air enters via duct 200 and process 
water via line 180. Flue gas exhausts through duct 202, carbon dioxide is 
vented via duct 182 and spent adsorbent and/or fluid bed fines are 
disposed of via transport system 188. Product hydrogen is recovered via 
line 184 for supply to units 8, 10 and 12, and may be supplemented, if 
required, by hydrogen supplied through line 204 from another source. 
If insufficient fuel is available from the residues supplied through line 
124, additional fuel such as coal or lignite may be supplied via transfer 
system 181 and control device 183. 
Referring now to FIG. 3, 802 is a as turbine-driven combustion air 
compressor and power generation unit including an air compressor 804, a 
hot gas expander 806 and an alternator 808. The machine is shown as a 
single fixed shaft machine but may also be a single split shaft machine or 
a machine having more than one shaft for air compression and hot gas 
expansion, and may also be split into a number of machines, for example 
one expander driving the air compressor and a separate expander driving 
the alternator. In all cases, the drive shaft to the compressor and/or 
alternator may incorporate gearboxes. 
810 is a pressurised fluid bed combustor/gas reformer having a distribution 
grid 812, a fluid bed 814 and a plurality of catalyst-filled reformer 
tubes 816 (for simplicity, only one is shown). 818 is a dust remover 
cyclone or set of cyclones or an alternate filtration device, 820 and 822 
are ash removal and pressure let-down devices and 824 is a flue gas waste 
heat boiler. 826 is a set of heat exchangers for feed gas preheating and 
hot synthesis gas cooling and 828 is a boiler. 830 is a shift reactor 
system, 832 is a carbon dioxide scrubbing system and 834 is a methanation 
system. 830, 832 and 834 are designed in accordance with well-established 
principles for the design of such units for hydrogen plants. 836 and 838 
are devices or systems for compressing and feeding, respectively, fuel for 
the fluidised bed combustor and crushed limestone and/or dolomite and/or 
inert material such as sand or alumina for the fluidised bed. 
Combustion air entering the system through pipeline 200 (FIG. 1) is drawn 
through duct 902 into the air compressor 804 in which it is compressed to 
between 5 and 40 Bar, and preferably to between 12 and 25 Bar. 
Intercooling during compression may be employed but after-cooling is not 
desirable. The compressed air is then passed via pipeline 904 to the base 
of the fluid bed combustor 810. 
Residues provided through line 124 (FIG. 1) connecting with line 922 are 
fed at a controlled rate and under pressure to the fluidised bed by 
feeding system 836. The feeding system may comprise a heater and pump in 
the case of fusible oil residues or, where the residues are infusible, 
e.g. as in the case of the char formed in the arrangement of FIG. 2 below, 
the system may comprise a water slurrying device and pumping system or a 
dry powder feed sysem such as manufactured by Petrocarb or some suitable 
alternate device. The fuel from the feeding device 836 passes via pipeline 
924 to the fluid bed combustor 810. Pipeline 924 may be a single line or 
more preferably a number of lines in parallel. The fuel is shown as fed 
into the fluid bed 814 above the distribution grid 812 but the feed point 
or points may be incorporated with the distribution grid 812. It is 
possible to feed fuel at many points in the fluid bed 814; however it is 
preferable to feed the fuel at least between the grid 812 and the lower 
manifold 944 (referred to below) and preferably close to the grid 812. 
Inert material such as sand or alumina to produce or augment the necessary 
fluid bed 814 and/or a sulphur adsorbent such as limestone or dolomite 
supplied by transport system 186 (FIG. 1) is fed via feed duct 926 to 
feeding device 838 which may be a water slurrying device and pumping 
system or a dry powder feed system or some suitable alternate device. The 
inert material and/or adsorbent is fed via pipeline 928 to fluidised bed 
814. The location of feed pipe 828 is not critical but ideally is above 
and adjacent to grid 812. It is possible to integrate feed devices 836 and 
838 to feed a mixture of fuel and inert and/or adsorbent or to slurry the 
inert and/or adsorbent in molten oil residues. 
The grid 812 may be a proprietory design or a heat resistant metal plate 
which may be in segmented form to allow for expansion and contraction, in 
which a series of bubble-cap type distributors or alternate gas 
distribution devices are fixed. Even distribution of air through the 
distributors and the bed requires a pressure drop across the distributors 
in the order of 0.05 to 0.10 Bar. 
For any given pressure, the range of air and flue gas velocities in the bed 
for a given average particle size is well-established and it is generally 
desirable to operate with a velocity in the lower range of available 
velocities in order to ensure minimum erosion of the reformer tubes 816, 
manifolds and connecting lines and risers 942, 944 and 946. From the 
velocity and the quantity of combustion air required to provide the 
necessary heating, the cross-sectional area of the bed can be determined. 
The depth of the bed may range from 2 to 14 meters or more. The preferred 
depth for vertical cylindrical and un-finned reformer tubes is 7 to 12 
meters with the tubes occupying the upper 6 to 11 meters of the bed. The 
fluid bed consists predominately of ash from the fuel and/or deliberately 
added inert material and/or dolomite or limestone adsorbent and in order 
for the bed height to be held constant, ash is drawn off from the bed 
through duct 814 to draw-off device 822. The draw-off may be constant or 
intermittent and may be based on the overflowing of the bed into the duct 
914 or it may be controlled by a bed level sensing device which may be 
designed to detect the pressure drop across the bed and by inference the 
bed height. The draw-off device draws off ash under pressure and may let 
it down through a system of lock-hoppers using well-established designs or 
through one or more rotary type valves or some suitable alternate device. 
De-pressurised ash and/or inert material and/or spent adsorbent is 
discharged via duct 916 to duct 188 (FIG. 1). 
The fluid bed may be of the simple up-flow type as illustrated or it may 
incorporate a spouting device with two separate and distinct upward 
velocity zones. 
It is preferred to operate the bed with an excess of air ranging between 
20% and 50% above stoichiometric. However it is possible to operate above 
50% excess air if the plant is required to produce a surplus of energy in 
the form of mechanical energy from the turbine 806 and/or in the form of 
steam. It is also possible to operate with a deficiency of air with the 
combustion of carbon monoxide and possibly hydrogen formed in the bed 
being carried out by adding more combustion air above the bed 814 or 
subsequent to the combustor 810. It is generally not desirable for 
metallurgical reasons to operate the bed such that there are transitions 
between an excess and deficiency of combustion air. 
The temperature of the fluid bed 814 is controlled at a fixed value and in 
the range of 750.degree. to 1100.degree. C. but more preferably within the 
range of 850.degree. to 1000.degree. C. generally by primary control of 
the fuel feed rate and secondary control of combustion air flow to 
minimise the likelihood of carbon formation in reformer tubes 816 or 
overheating of the connecting tube headers and risers 938, 944, 946. This 
temperature range also allows the use of well-established alloys such as 
Incalloy and HK40 for tubes support etc and also the effective retention 
of sulphur by limestone or dolomite adsorbent if sulphur removal is 
required. 
Hot fuel gases leave the combustor 810 through the duct 906. These hot 
gases also contain fine dust in suspension and the mixture passes through 
the dust removal system 818 which may consist of cyclones and/or 
filtration systems but generally consists of primary and secondary cyclone 
systems which ensure the removal of the greater part of the entrained 
dust. The maximum size of dust particle leaving the secondary cyclone 
should be about 10 micron so as not to unduly interfere with the operation 
of the hot gas expander 806. Ash is removed from separator 818 via pipe 
918 to the ash discharge device 820 which may be a system of lock-hoppers 
and/or rotary valves or similar alternate device. De-pressurised ash is 
removed via pipeline 920. Pressurised and cleaned hot gas from separator 
818 passes via duct 908 to the expander turbine 806 in which the exhaust 
gases are let down to substantially atmospheric pressure and then pass 
through duct 910 to a waste heat boiler 824 and thence via duct 912 to 
atmosphere. Suitably pressurised feed water is admitted via pipeline 960 
to the boiler 824 which may incorporate a preheater and superheater and 
the resultant produced steam leaves via pipeline 962. 
Methane-containing gas from line 178 is supplied via line 930 to the heat 
exchange system 826 where a controlled amount of steam is added from 
pipeline 966. In 826, the methane/steam mixture is heated to about 
500.degree. C. before passing via pipeline 936 into the pressurised 
combustor 810 where it is distributed by the header system 938 and a 
series of connecting tubes 940 to catalyst filled reformer tubes 816 (one 
only is shown). These tubes may be of current state of the art design i.e. 
vertical, cylindrical, plain tubes packed with a suitable proprietary 
steam reformer catalyst or they may be finned tubes or abnormally shaped 
tubes which would operate satisfactorily due to the reduced stresses on 
the tubes. The methane and any heavier hydrocarbons associated with the 
methane are substantially reformed in the tubes to hydrogen and carbon 
monoxide and exit through connecting stubs 942, bottom collection header 
944 and thence upwardly through riser 946 and through the shell of the 
combustor 810 to the heat exchange system 826. The riser 946 and the feed 
line 936 together with the headers 938 and 944 incorporate suitable 
arrangements for differential expansion and contraction due to the 
temperature variations during start-up operation and shut-down of the 
combustor. The tubes are suitably supported using for example a high 
temperature metal support grid above the tubes supported from the walls of 
the combustor 810 using tie rods to support the upper tube header 938, the 
inlet pipe 936 and the riser 946. 
In heat exchange system 826, the hot reformate is cooled to about 
500.degree. to 650.degree. C. and is then transferred via line 948 to 
shift reactor system 30 where it is mixed with water and/or steam to 
further cool the gas to about 360.degree. to 380.degree. C., and then 
passed over a high temperature shift catalyst. It is then mixed with 
additional water and/or steam to further cool the gas to about 210.degree. 
C. after which it is passed over a low temperature shift catalyst to 
convert the bulk of the carbon monoxide by reaction with water to carbon 
dioxide and hydrogen. The product gas passes via line 952 to the cooling 
and carbon dioxide scrubbing system 832 where the reformed and 
shift-reacted gas is cooled and scrubbed to remove carbon dioxide. The 
heat content of the entering gases is generally sufficient to provide the 
necessary heat for stripping the carbon dioxide from the scrubbing liquid 
used for the process, namely hot carbonate solution. Carbon dioxide is 
withdrawn via line 954 and condensate via line 980. 
Gas recovered from the cooling and scrubbing unit 832 is passed via line 
956 to methanation unit 834 in which any remaining carbon monoxide is 
converted to methane by known means. The unit may incorporate a separately 
fired heater or may (by means not shown) employ heat from the hot gas 
streams in lines 910 and/or 958. The gas is then further cooled in the 
methanation unit and condensate is recovered via line 982. Product 
hydrogen is recovered via line 958 which connects with line 184 (FIG. 1) 
for passing the hydrogen to the hydrogenation units 8, 10 and 12. 
The steam for the steam reforming may be provided from line 962. If further 
steam is required for the steam reforming and/or for other uses e.g. power 
generation, it may be provided from line 978 from ancillary boiler 828. 
Feed water at a suitable pressure is supplied to this boiler via line 964 
and water is circulated via lines 972, 974 and 976 through the fluid bed 
814 where it is heated to produce superheated steam or a mixture of water 
and steam. Steam generated in this boiler is recovered via line 978. 
If desired, feed water for boiler 828, and also boiler 824, may be 
preheated in the cooling and carbon dioxide scrubbing unit 832. 
By means of the arrangement illustrated in FIG. 3, hydrogen may be produced 
at about 20 Bar with a purity of up to 98%. If the pressure is increased 
to 40 Bar, the product gas will contain about 85 to 90% hydrogen. However, 
as the purity of the hydrogen is not critical for the hydrogenation 
reactions in hydrogenation units 8, 10 and 12, such lower purity is 
acceptable and operating at this higher pressure reduces compressor and 
compression costs. 
With the combustor 810 running at its preferred pressure range, expander 
turbine 806 will produce energy in excess of that required to power 
compressor 804. Recovery of the surplus energy plus speed control of the 
turbine 806 may be achieved by generating power in alternator 808 which 
supplies electric power via cable 968/212. (FIG. 1) 
A coal-based synthetic oil plant incorporating the process of the invention 
is illustrated in FIG. 2 in which reference numeral 402 is a solvent 
extraction unit designed to operate a sub-critical solvent extraction 
process, a super-critical solvent extraction process or a solvent 
contacting process combined with hydrogenation, i.e. H-coal process. 4 is 
a separation unit for separating the product from 402 into a liquid 
extract (possibly in solution with a solvent ) which is suitable for 
hydrogenation, and a solid residue comprising mineral matter and coal 
char. It will be understood that units 402 and 404 could be replaced by a 
pyrolysis or carbonisation plant. As in FIG. 1, 8, 10 and 12 are 
hydrogenation units, 14 is a unit for the atmospheric distillation of 
hydrogenated material recovered from units 8, 10 and 12, and 16 is a unit 
for the vacuum distillation of residue from unit 14. Similarly, units 20, 
22 and 24 are, respectively, a gas purification plant, gas separation 
plant and hydrogen plant. 
It will be understood that, as in the case of the plant illustrated in FIG. 
1, other units such as catalytic reformers and catalytic crackers may also 
be included, and that the units illustrated incorporate vessels, pumps, 
heaters etc in well known manner. Variations in the flow arrangements are 
possible and the lines shown may represent a single pipeline or a 
plurality of pipelines arranged in parallel. 
Suitably prepared coal is fed to solvent extraction unit 402 by means of 
transportation system 502 and solvent is provided through line 508 and/or 
506. Where hydrogenation is required, hydrogen may be provided via line 
512. Any surplus gas is vented via lines 510 and 516 to line 552. 
The mixture of dissolved coal, solvent, undissolved char and mineral matter 
is passed by means of transport system 504, which may be a lock hopper, 
pipeline or suitable alternate device, to separation unit 404. In this 
unit, generally by either pressure reduction and/or heating, additional 
gas is liberated and removed via line 514. Surplus solvent is recovered 
via line 506 and recycled to solvent extraction unit 402, and coal 
liquids, possibly still containing some solvent, are recovered via line 
522 and passed via line 546 to hydrogenation unit 12. 
Hydrogenated material produced in unit 12 exits via line 550 and is passed 
via line 554 to distillation unit 14. All the lines shown in association 
with units 8, 10, 12, 14 and 16 have the same function as the 
corresponding lines of FIG. 1 having the same last two digits in the 
reference numeral, with the exception that distillates in lines 532, 533, 
543 and/or 568 may be recycled to solvent extraction unit 402 via line 
508. Further, in some cases residue from distillation unit 14 may be 
recovered through line 561 as product, and the vacuum distillation unit 16 
may be omitted. 
Gas from hydrogenation units 8, 10 and 12, distillation unit 14, solvent 
extraction unit 402 and/or separation unit 404 is collected in line 552 
and passed to gas purification unit 20, gas separation unit 22 and 
hydrogen plant 24, all of which operate in the same manner as in the 
arrangement described with reference to FIG. 1. 
However, in this plant, the hydrogen production unit 24 is fired with coal 
residues from separation unit 404 and supplied through transport system 
524. If desired, part of the residue may be disposed of via control device 
525 and transport system 527. Likewise, if insufficient residue from 
separation unit 404 is available to fuel hydrogen plant 24, additional 
fuel may be provided via transport system 581 and control device 583. 
Thus, in this plant, feed coal provided through transport system 502 is 
treated in unit 402 and separated in unit 404 into a liquid product and a 
residue. The liquid product is hydrogenated in 12, hydrogenated material 
is fractionated in 14 and 16 and liquid streams recovered from these units 
are further hydrogenated in hydrogenation units 8 and 10. Hydrogenated 
material from 8 and 10 is recycled to distillation units 14 and 16 and the 
desired products are recovered through lines 532, 533, 543 and/or 568. The 
gas recovered from units 8, 10, 12, 14, 402 and/or 404 is steam reformed 
in hydrogen plant 24 to produce the required hydrogen for plants 8, 10 and 
12. The fuel-containing residues from separation unit 404 provide the fuel 
for the fluidised bed combustor of hydrogen plant 24 which in detail is as 
described above with reference to FIG. 3 but with the fuel feed supplied 
to line 922 being coal residues provided via transport system 524. 
EXAMPLE 1 
Using the plant described and illustrated in FIGS. 1 and 3, 26,178 
tonnes/day of a blended atmospheric residue oil derived from middle east 
crude oils is supplied through line 110 to vacuum distillation unit 4 
where it is separated into a distillate all of which is hydrogenated in 
hydrogenation unit 10 and a residue which is recovered through line 
119/124. Liquid hydrogenate from 10 is distilled in atmospheric 
distillation unit 14 to yield light naphtha, heavy naphtha, kerosine and 
diesel cuts, and a residue which is distilled in vacuum distillation unit 
16 to produce a distillate which is recycled to hydrogenation unit 10 and 
a residue. Gas from the hydrogenation is recovered in line 142 and, after 
treatment in 20 and 22 to remove ammonia and H.sub.2 S, is passed to 
hydrogen generation plant 24 from which the hydrogen product is returned 
in lines 184 and 208 to the hydrogenation unit 10. Units 2, 6, 8 and 12 
are not used. 
The process details are as follows. 
______________________________________ 
Vacuum gas oil recovered in line 114 
______________________________________ 
A Flow rate 16492 tonnes/day 
B Boiling range: IBP 347.degree. C. 
C 50% 450.degree. C. 
D EP 505.degree. C. 
E Sulphur content 1.8% w/w 
F Nitrogen content 0.1% w/w 
______________________________________ 
______________________________________ 
Vacuum residue recovered in line 119/124 
______________________________________ 
G Flow rate 9686 tonnes/day 
H A.P.I. gravity 5.9 
H Sulphur content 4.5 % (wt/wt) 
K Ni and V content 130 ppm 
______________________________________ 
______________________________________ 
Composition of gas in line 142 (weight % of feed flow in line 
______________________________________ 
110) 
H.sub.2 S and ammonia 2.04% 
M Methane, ethane, propane, butane and light 
naptha 5.63% 
______________________________________ 
______________________________________ 
Hydrogen in line 184/208 
______________________________________ 
N Flow rate 419 tonnes/day 
______________________________________ 
______________________________________ 
Feed to Hydrogen Plant 
______________________________________ 
P Vacuum residue in line 124: 
910 tonnes/day 
R Dolomite 350 tonnes/day 
S Steam at 25 Bar 7210 tonnes/day 
______________________________________ 
______________________________________ 
Product streams recovered from Atmospheric distillation unit 
______________________________________ 
14 
T Light naphtha 
434 tonnes/day 
V Heavy naphtha 
1145 tonnes/day 
W Kerosine 6465 tonnes/day 
X Diesel 7603 tonnes/day 
______________________________________ 
EXAMPLE 2 
Using the plant described and illustrated in FIGS. 2 and 3, coal in 
powdered form and provided through line 502 is solvent extracted by 
supercritical solvent extraction techniques in 402 and the coal liquids 
thereby produced are separated in 404 and hydrogenated in hydrogenation 
unit 12. Gas from 12 is recovered in line 148/552 and passed to the 
hydrogen plant 24 and hydrogen generated from this gas is returned to 12 
in line 584/606. Liquids from the hydrogenation plant are passed in line 
550/554 for atmospheric distillation in 14 to produce two liquid products, 
recovered through lines 556 and 560 respectively. Gas from 402 and 14 
joins the gas from hydrogenation unit 12 in line 552 for passing to 
hydrogen plant 24. Char recovered from separation unit 404 in line 524 is 
passed to the hydrogen plant 24 as fuel for the fluid bed combustor. Units 
8, 10 and 16 are not used. In this Example, about 75 tonnes/day of gas is 
purged through line 176. 
The process details are as follows. 
______________________________________ 
AA Feed coal in line 502: Flow rate (ash 
and moisture free) 23,650 
tonnes/day 
CC Char in line 524, production rate (ash 
and moisture free) 13,782 
tonnes/day 
DD Total hydrocarbon gas in line 552 including 
44.1 tonnes/day H.sub.2 1,203.6 
tonnes/day 
EE Rate of consumption of char in hydrogen 
plant 24 1,692 
tonnes/day 
FF Net surplus char available as product via 
line 527 (including 40% ash) 
21,278 
tonnes/day 
Liquid fractions available as product from 
distillation unit 14: 
GG in line 556/532 IBP 200.degree. C. fractions 
2,784 
tonnes/day 
HH in line 560 200.degree. C. + fraction 
4,923 
tonnes/day 
______________________________________ 
The following illustrates the operating details of steam reformer/fluid bed 
combustor described and illustrated in FIG. 3 and employed in hydrogen 
plant 24 of Example 2. 
__________________________________________________________________________ 
JJ Char in line 924: calorific value 
5,054 
Kcal/kg 
KK ash content 40% 
LL feed rate 1,692 
tonnes/day 
MM average particle size 0.6 mm 
Feed gas in line 930 (C/H ratio is 
3.8/1 wt/wt) 
NN pressure 25 Bar 
PP feed rate 1128.3 
tonnes/day 
RR Number of combustors 810 
4 
SS Inside diameter of each combustor 
5.8 meters 
TT Number of reformer tubes in each combustor 
500 
VV Outside diameter of reformer tube 
100 mm 
WW Length of reformer tube 
9.15 meters 
XX Depth of fluid bed 11.0 meters 
ZZ Air flow in line 904: flow rate 
565,000 
Nm.sup.3 /hr 
AAA pressure 20 Bar 
BBB Steam/methane ratio in 936 4.0 
CCC Hydrogen product in line 958: flow rate 
441 tonnes/day 
(as 100% hydrogen) 
DDD purity (minimum) 95% hydrogen 
EEE pressure 18 bar 
FFF Adsorbed power of air compressor 804 
95,000 
Kw 
GGG Power generated in turbine 806 
140,000 
Kw 
HHH Net power available from alternator 808 
in 
line 968 45,000 
Kw 
JJJ Power required for hydrogenation plant in 
Example 1 45,000 
Kw(approx) 
__________________________________________________________________________ 
This steam reformer/fluid bed combustor is employed in the arrangement of 
FIG. 1, by replacing char feed by a supply of vacuum distillation residue 
provided through line 119/124 (FIG. 1). 
In both Example 1 and Example 2, water is injected into the fluidised bed. 
In Example 2, it is provided by feeding the char as 50/50 wt/wt slurry in 
water. In Example 1, an equivalent amount of water is added with the 
dolomite. Where the fluidised bed is pressurised, injecting water 
increases the power available from the flue gas and is a valuable means 
for disposing of waste water from processes involved in the extraction of 
the residue used to fuel the bed e.g. colliery waste water or refinery 
oily water waste.