Production of methanol

In the production of methanol from a methane-containing gaseous feedstock such as natural gas by steam reforming the gas and treating the reformate to produce methanol by inter-reaction of the hydrogen and oxides of carbon in the reformate, the use of part of the feedstock to fire the reformer is avoided by immersing the reformer reactor tubes in a fluidized bed heated by the combustion of a low grade, solid, fossil-based fuel such as coal, lignite, oil shale or asphaltic residues from oil refining. By pressurizing the fluidized bed, all the power requirements of the process can be obtained by expansion of the flue gas, which can also provide the CO.sub.2 balance for the methanol synthesis, and a compressor can be omitted.

This invention relates to the production of methanol from a gas, such as 
natural gas, which contains methane as a principal component. 
Natural gas is frequently found in locations which are distant from 
industrialised communities and transport is a problem. At present, 
transport is effected by pumping the gas through a pipeline or by 
liquefying the gas and then transporting it by pipeline or by tanker. An 
attractive alternative, however, would be to convert the gas to methanol 
since this is easier and safer to handle and transport. 
Another factor which favours examination of the practicality of converting 
natural gas to methanol is that whereas natural gas cannot readily be 
employed as a transport fuel, the use of methanol for this purpose is well 
recognised. Thus, providing methanol from natural gas, of which there are 
large reserves, would ease the current total dependence of the transport 
industry on the diminishing reserves of oil. 
It is known that natural gas may be converted to methanol by steam 
reforming the natural gas to produce a gaseous reformate containing 
hydrogen and oxides of carbon and forming methanol from this reformate by 
reaction of the oxides of carbon with the hydrogen, e.g. in accordance 
with the following simplified equations: 
Reforming 
EQU CH.sub.4 +H.sub.2 O.revreaction.CO+3H.sub.2 ( 1) 
EQU CH.sub.4 +2H.sub.2 O.revreaction.CO.sub.2 +4H.sub.2 ( 2) 
Methanol Production 
EQU CO+2H.sub.2 .revreaction.CH.sub.3 OH (3) 
EQU CO.sub.2 +3H.sub.2 .revreaction.CH.sub.3 OH+H.sub.2 O (4) 
However, the reforming reaction is endothermic and is conducted at elevated 
temperatures, e.g. 800.degree. C. to 900.degree. C., and the methanol 
production step requires a higher pressure than the steam reforming step. 
Thus both heat for the reforming and energy for gas compression are 
required. In existing processes, these heat and energy requirements are 
supplied by combustion of a portion of the natural gas. This is because 
the use of a premium fuel such as natural gas is regarded as essential for 
the achievement of the critical degree of temperature control necessary 
for the practical operation of the reforming process and as the natural 
gas is already available as the chemical feedstock it is the natural first 
choice for the fuel. The overall efficiency of the production of the 
methanol from the natural gas, therefore, cannot readily exceed about 68% 
and this shortcoming has been the major factor in preventing development 
of this process as an alternative to pipelining the natural gas or 
transporting it as LNG and using it ultimately as a gaseous fuel. It has 
also been recognised for over a decade as constituting a serious 
disadvantage of the natural gas route to methanol and thus although the 
potential benefits of providing methanol from natural gas have been 
recognised for many years and have received particular attention and 
publicity since the late 1960s, in fact the process has in general only 
been employed for the production of methanol for chemical purposes. 
Many proposals have been made for increasing the efficiency of the 
utilisation of the gas in the production of methanol. However, even though 
since 1973, or even earlier, the premium value of natural gas has been 
clearly established, and for upwards of a decade major studies have been 
undertaken to determine the optimum manner in which to develop gas 
resources in the context of known regional reserves of gas and other 
fuels, and in specific instances to consider the production of methanol 
from gas and proximate solid fuel reserves, to date these proposals have 
been directed at improvements in heat recovery despite the limited 
increase in efficiency that can be attained thereby; and although for many 
years there has been a major incentive to use low grade fuels in place of 
natural gas wherever possible, hitherto no practical proposal has been 
published for a process for producing methanol from natural gas wherein an 
alternate lower grade fuel is used to heat and power the plant. 
The present invention provides a practicable route to methanol from natural 
gas or other methane-containing gas wherein the heat and energy 
requirements are supplied by the combustion of an alternate low grade 
fuel. 
According to the present invention, there is provided a method of producing 
methanol from a methane-containing gas by the steps of (a) steam reforming 
said gas at elevated temperature to form a reformate containing hydrogen 
and oxides of carbon and (b) forming methanol from the reformate by 
reaction of the hydrogen with the oxides of carbon, and wherein the step 
of reforming the gas is effected in the presence of a catalyst within a 
reactor vessel and heat for the reaction is provided by at least partly 
immersing said reactor vessel in a heated fluidised bed of finely divided 
solid and at least a part, and preferably all, of the heat is provided by 
the combustion of a solid, fossil-based fuel. 
For economic reasons it is preferred that the gas contains at least 40% by 
volume of methane and preferably more but gases containing less methane 
can be used. The balance, if any, will depend on the source of the gas but 
may include, for example, one or a mixture of the following: inert gases 
e.g. argon, nitrogen and helium, hydrogen, oxides of carbon, and 
hydrocarbons containing more than one carbon atom, e.g. ethane, propane, 
ethylene and propylene. Examples of such gases are natural gas, gas 
associated with crude oil recovery (known as associated gas), mine 
drainage gas, gas obtained from the biodegradation of organic matter, gas 
produced during the processing of crude oil or as a by-product in the 
production of synthetic oil, e.g. from coal, and purge gas from the 
production of methanol from methane-containing synthesis gas produced by 
the gasification of a solid fuel or oil. However, the process is 
particularly suitable for the processing of natural gas and mine drainage 
gas. Typically, such gases contain at least 70% methane, by volume. 
By a solid fossil-based fuel is meant a fossil-based fuel which is normally 
solid at room temperatures. Examples are coal, lignite, peat, oil shale 
and oil-derived solid residues such as, for example, vacuum distillation 
residues, deasphalter residues and petroleum derived coke. 
In the steam reforming step, the methane is reacted with steam in the 
presence of a catalyst to form hydrogen and oxides of carbon, e.g. in 
accordance with idealised equations (1) and (2) above. The conditions and 
catalysts appropriate for this steam reforming step are well known and 
generally include temperatures in the range 700.degree.-1000.degree. C., 
preferably 800.degree.-900.degree. C., and pressures in the range 100-600 
psig, preferably 200 to 300 psig. The steam: methane ratio is generally 
1:1-6:1, preferably about 4:1. 
The process is preferably operated to achieve as high a conversion as 
possible and conversions of a high as 90% or higher, e.g. 95% to 98% may 
be achieved. 
The reaction is preferably effected by passing the gaseous mixture through 
one or more heated tubes containing the catalyst. The manner of providing 
the heat is described below. 
The reformate recovered from the reformer, and which will contain hydrogen, 
oxides of carbon, inert gases and unreacted methane, is then subjected to 
conditions under which the hydrogen reacts with the oxides of carbon e.g. 
in accordance with idealised equations (3) and (4). The conditions and 
apparatus appropriate for such methanol synthesis are well known and 
generally include temperatures in the range of 200.degree. to 300.degree. 
C. and pressures in the range 40-300 bar, preferably 65-100 bar. 
It is preferred for the feed to the methanol synthesis reaction to include 
added carbon dioxide since this increases the amount of methanol that can 
be produced per mole of methane. The carbon dioxide may be added to the 
reformer feedstock but preferably it is added to the reformate prior to 
the methanol synthesis reaction. In the absence of carbon dioxide, the 
overall reaction for the production of methanol from methane by steam 
reforming can be represented as 
EQU CH.sub.4 +H.sub.2 O.revreaction.CH.sub.3 OH+H.sub.2. 
In the presence of carbon dioxide, however, the reaction becomes 
EQU CH.sub.4 +1/3CO.sub.2 +2/3H.sub.2 O.revreaction.4/3CH.sub.3 OH. 
The methanol synthesis may be effected in any suitable apparatus. As the 
temperature preferred for the methanol synthesis is substantially below 
that of the reforming step, in general it is necessary to cool the 
reformate prior to contacting it with the methanol synthesis gas. However, 
in accordance with Le Chatelier's principle, methanol production is 
favoured by high pressures and the pressures preferred for the methanol 
synthesis will normally be higher than those preferred for the reforming 
step in which case it will be necessary to compress the reformate. 
In general, the art has found it expedient to operate the methanol 
synthesis at a relatively low conversion, e.g. 10 to 20%, and it is 
conventional to recycle the unreacted gases after treatment to remove the 
methanol e.g. by condensation and/or scrubbing. Because of the pressure 
drop through the reactor, it is necessary to recompress the recycle stream 
and because of the low conversion this recycle stream forms a large 
proportion of the total feed to the methanol synthesis reactor. To prevent 
the build-up of non-reactants in the reactor, a purge stream of 
appropriate size is usually bled from the recycle stream and may be 
employed as a fuel or a process gas. 
The crude methanol recovered from the product stream from the methanol 
synthesis reactor may be purified in conventional manner. 
In accordance with the invention, the heat required for the reforming step 
is provided by at least partly immersing the reformer reactor vessel, 
which may comprise a single catalyst-packed tube but more usually a 
plurality of catalyst packed tubes arranged in parallel with reference to 
the flow of process gas, in a fluidised bed of finely divided solid which 
is heated by the combustion of a solid fossil-based fuel. Such fluidised 
bed combustors have been known for decades and the nature and operation 
thereof are well publicised and understood. 
An important aspect of the present invention, however, is the realisation 
that by employing a fluidised bed combustor to provide the heat to the 
reformer tubes, the problems hitherto associated with using a solid fuel, 
and in particular the problems of controlling the distribution of heat to 
the reformer reactor walls and avoiding the build up of deposits on the 
reactor walls, can be overcome and it is now possible to avoid using the 
valuable reformer feedstock for this purpose. In fact, not only can a 
sufficiently uniform heat transfer to the walls of the reactor tube be 
achieved, but temperature control can be improved over that obtainable 
using the premium gaseous fuel in a conventional radiant heat furnace. 
Other advantages also accrue from the use of a fluidised bed combustor; in 
particular, 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 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 has recently been proposed to form reformer reactor tube walls of 
ceramic materials in order to permit higher operating temperatures. Having 
the fluidised bed under superatmospheric pressure permits the placing of 
such tubes under compression, in which mode they offer better performance. 
Yet another advantage of having the fluidised bed under pressure is that 
the furnaces will be basically totally enclosed and the design is 
therefore more suited to flame proof or generally hazardous areas than a 
conventional furnace. 
The fluidised bed furnace in addition to being more compact will also 
contain much less refractory than, and need not have flat wall refractory 
as in, a conventional furnace. The fluidised bed furnace will therefore be 
more suited to transportation or installation in moving and floating 
plants such as may be required for the treatment of off-shore gases. 
A very important advantage of having the fluidised bed at superatmospheric 
pressure is that energy in the flue gas from the combustor may be 
recovered by expanding the gas through an expansion engine such as a 
turbine which may be employed for example to drive a generator and/or to 
provide part or all of the compression energy requirements of the process, 
e.g. to compress the gaseous oxidant e.g. air to the combustor and/or to 
compress the reformate to the higher pressure required for the methanol 
production step and/or to recompress any recycle streams in the system 
e.g. unreacted or partially reacted reformate being recycled to the 
methanol production step. The extraction of energy directly in this 
fashion is more efficient than the conventional raising of high pressure 
superheated steam and the use of this steam in driving the process 
compressors or generating electricity. 
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 of 20% 
to 50% in 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. 
By suitable choice of the pressure in the combustor and the amount of 
excess air supplied to the combustor, all the compression energy 
requirements of the process can be satisfied from the energy generated by 
the turbine. 
As indicated above, improvements in the possible yield of methanol per mole 
of methane consumed may be achieved by injecting carbon dioxide into the 
gas mixture before reforming and/or before the methanol synthesis step. In 
accordance with another preferred embodiment of the invention, at least a 
part and preferably all of the carbon dioxide is provided from the flue 
gas from the fluidised bed combustor. The recovery of the carbon dioxide 
from the flue gas is facilitated by operating the combustor at 
superatmospheric pressure. Suitably, a part of the flue gas, generally up 
to about 25% thereof, is removed from the main portion of the flue gas, 
treated to recover carbon dioxide therefrom, and thereafter recombined 
with the main portion of the flue gas prior to expansion of the latter 
through a turbine to recover energy therefrom. In one embodiment, the 
carbon dioxide is recovered by washing in a scrubber and the scrubbed gas 
is then preferably reheated, e.g. by heat exchange with the gas being fed 
to the scrubber, prior to recombination with the main portion of the flue 
gas. 
A further important aspect of the invention is that operating the fluidised 
bed combustor under superatmospheric pressure offers the possibility of 
increasing the pressure of the reforming reaction within the reactor tubes 
towards that preferred for the methanol synthesis. While such increase in 
reforming pressure will tend to reduce the conversion of the methane and 
thus increase the amount of unreacted methane in the reformate, this 
increase can be dealt with by diverting a part of the recycle stream of 
the methanol synthesis reaction back to the reformer. This will increase 
the heat requirement for the reforming step but since the fuel for the 
reformer is not provided from the feedstock, the overall feedstock 
conversion efficiency will not be affected. Such operation has the 
advantage of permitting a substantial reduction of the equipment required 
for cooling the hot reformate and a preferred arrangement, in which the 
reformate is produced at substantially the same pressure as the methanol 
synthesis recycle stream whereby the two may be combined for feeding to 
the methanol synthesis reactor, confines the primary compressor 
requirement to the recirculation of the methanol synthesis recycle stream 
thus permitting the use of large volume flow low pressure ratio 
compression equipment and leading to a substantial reduction in the 
capital cost of the plant as well as simplification of the process. A 
single compressor may be used for the recirculation and reformate 
pressurisation, if desired. 
Where the fuel for the fluidised bed is an ash-containing solid, such as 
coal, coke, lignite, oil shale or peat, 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 as 
or 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 liquid 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, minerals such as 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 6 mm, and an average particle size in the range of 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. 
In many locations, natural gas is associated with fields of solid fuel or 
with the recovery of crude oil and the process of the invention therefore 
permits the establishment of the process at a location where both the 
feedstock for the methanol synthesis and the fuel for providing the heat 
and power requirements of the process are available locally. Methane gas 
can also be recovered from some coal fields by established coal seam 
drainage techniques. In this case, any mining operations allied to the 
seam drainage can also be allied to the operations for obtaining the fuel. 
Preferably the coal is separated into a low ash fraction and a high ash 
fraction ("high" and "low" meaning higher and lower, respectively, than 
the average ash content of the coal), and the high ash fraction is 
employed as the fuel for the process. The low ash fraction may then be 
employed e.g. for metallurgical processes. 
In another aspect of the invention, both the feedstock and the fuel may be 
provided from the by-products of oil refining. For example, the feedstock 
may be supplied from the gas recovered from crude oil separation and 
distillation and/or such operations as hydrogenation, hydrotreating, 
hydrocracking and cat cracking, and the fuel may be provided from the 
residue of atmospheric and/or vacuum distillation and/or the residues from 
deasphalting and/or demineralisation operations. 
The invention is now described in greater detail with reference to 
preferred embodiments thereof and with the aid of the accompanying 
drawings in which 
FIG. 1 is a block flow diagram of one embodiment of the process of the 
invention 
FIG. 2 illustrates how the arrangement of FIG. 1 may be simplified when the 
reformate is produced at substantially the same pressure as the recycle 
gas from the methanol synthesis reaction, and 
FIG. 3 illustrates the application of the invention to a coal mining and 
mine gas drainage complex.

Referring to FIG. 1, 2 is a gas turbine-driven compressor and power 
generation unit comprising an air compressor 4 and hot gas expander 6, a 
synthesis gas compressor 8, a recycle compressor 10 and an alternator 12. 
The machine is shown in the diagram as a single fixed shaft machine but 
may in practice 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 machine expanding 
sufficient gas to drive a compressor and a separate machine or machines 
with gas expanders to drive the compressors 8 and 10 and the alternator 
12, and in addition each compressor may be divided into separate 
compressors. In all cases the drive shafts to the compressors and 
alternator may incorporate gearboxes. 
14 is a pressurised fluid bed combustor/gas reformer which includes a 
distribution grid 16 a fluid bed 18 and a catalyst filled reformer tube or 
tubes 20. 22 is a dust removal cyclone or set of cyclones or alternate 
filtration device; 24 and 26 are ash removal and pressure let-down devices 
and 28 and 29 are startup combustors. 30 is a flue gas waste heat boiler. 
32 is a gas/gas heat exchanger. 34 is a carbon dioxide scrubbing system 
and 36 is a carbon dioxide purification and compression system. 38 is a 
hot synthesis gas cooling and feed gas preheating set of heat exchangers 
incorporating a waste heat boiler or boilers and 40 is a zinc oxide 
purification system. 
42 is a methanol synthesis reactor which for the purposes of this 
description may be a catalyst filled tubular reactor in which the tubes 
are the heating tubes in a waste heat boiler. 44 is a gas/gas heat 
exchanger and reactor gas cooling and separation system and 46 is a crude 
methanol flash tank, treatment and distillation unit. 
It should be noted that items 8, 10, 38, 40, 42, 44 and 46 comprise a hot 
synthesis gas fed methanol plant which may be designed in accordance with 
the well established principles for the design of such plants. 
48 is a device or system for compressing and feeding crushed limestone 
and/or dolomite and/or inert material such as sand or alumina. 50 is a 
device or system for compressing and feeding a solid fossil fuel or fuels. 
Air is drawn through duct (102) into the air compressor 4 in which it is 
compressed to between 5 and 40 bar, but preferably in the range between 12 
and 25 bar. Inter-cooling during compression may be employed but 
after-cooling is not desirable. Compressed air passes via pipe 104 thence 
pipe 108, valve 110, pipe 112 and pipe 114 to the base of the fluid bed 
combustor 14. 
For start-up, by control of valves 110 and 232 all or part of the 
compressed air in 104 is diverted to pipe 106 and thence via pipe 230 and 
valve 232 to combustor 29. Fuel at a suitable pressure is supplied via 
control valve 636 and pipe 234 to combustor 29 and hot combustor gases at 
a suitable controlled temperature pass via pipe 238 to pipe 124 and thence 
to the hot gas expander 6. In this manner with a suitable starter motor 
for the gas turbine 2 and with the compressors 8 and 10 and alternator 12 
unloaded, the turbine may be started as a conventional gas turbine. 
Air to start up combustor 28 passes through pipe 240 and is controlled by 
valve 242 and fuel at a suitable pressure in pipe 244 is controlled by a 
valve 246 to produce hot combusted gases in pipe 248 which pass via pipe 
114 to the base of the fluid bed combustor and by control of valves 110 
and 242 the partially combusted air entering the base of the fluid bed 
combustor 14 may be raised at a controlled rate up to a final start-up 
temperature in the range of 700.degree. C. to 900.degree. C. Alternative 
designs incorporating start up burners in the fluid bed 18 or in the 
support grid 16 are also possible. 
From conveying systems 224, a solid fossil fuel, e.g. coal, lignite or 
solid oil residue, is fed forward at a controlled rate and under pressure 
by the feeding system 50. The feeding device may be a heater and pump in 
the case of oil residues or it may be a water slurrying device and pumping 
system for coal and lignite or it may be a dry powder feed system such as 
manufactured by Petrocarb or some suitable alternate device. The fuel from 
the feeding device 50 passes via pipeline 226 to the fluid fed combustor 
14. The pipeline 226 may be a single line or more preferably a number of 
lines in parallel. The fuel is shown as fed into the fluid bed 18 above 
the distribution grid 16 but the feed point or points may be incorporated 
with the support grid 16. It is possible to feed fuel at many points in 
the fluid bed 18; however it is preferable to feed the fuel at least 
between the grid 16 and the lower manifold 180 and preferably close to the 
grid 16. 
Inert material such as sand or alumina to produce the necessary fluid bed 
18 should the fuel be relatively ash free, and/or sulphur adsorbent such 
as limestone or dolomite is fed via the feed duct 220 to the feeding 
device 48 which may be a water slurrying device and pumping system or it 
may be a dry powder feed synthesis such as manufactured by Petrocarb or 
some suitable alternate device. The inert material and/or adsorbent is fed 
via pipeline 222 to the fluidised bed 18 in the combustor 14. The location 
of the feed pipe 222 is not critical but an ideal location is above and 
adjacent to the grid 16. It is also possible to integrate the feeding 
devices 48 and 50 to feed a mixture of fuel and inert and/or adsorbent 
material or to slurry the inert or adsorbent material in molten oil 
residues. 
The grid 16 may be one of a number of proprietary designs or it may be 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. In order to 
ensure even distribution of air through the distributors and the bed, the 
distributors should have a pressure drop in the order of 0.05 to 0.10 bar. 
The cross-sectional area of the bed is determined by the fluidising 
velocity and the volume of air required for the desired heat transfer. For 
any given pressure, the range of fluidising 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 20 and 
manifolds and connecting lines and risers 180,178 and 182 respectively. 
The depth of the bed 18 may range from as low as 2 meters to 14 meters and 
possibly more. However the preferred depth for vertical cylindrical and 
un-finned reformer tubes is in the range of 7 to 12 meters with the tubes 
occupying the upper 6 to 11 meters. The fluid bed consists predominantly 
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 150 to the draw-off 
device 26. The draw-off may be constant or intermittent and may be based 
on the overflowing of the bed into the duct 150 or it may be controlled by 
a bed level sensing device which in turn may be a device detecting the 
pressure drop across the bed and by inference the bed height. The draw-off 
device 26 draws off ash under pressure and may let it down through a 
system of lock-hoppers using well established designs or it may be let 
down through one or more rotary type valves or some alternate suitable 
device. Depressurised ash and/or inert material and/or spent adsorbent is 
discharged via duct 152. 
The fluid bed is preferably of the simple up-flow type as shown in the 
diagram but it may also be a bed as described in the literature 
incorporating a spouting device and 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. It is possible however to operate the 
bed at above 50% excess air if the methanol plant is required to produce a 
surplus of energy in the form of mechanical energy from the turbine 2 
and/or in the form of steam. It is also possible to operate the fluid bed 
with a deficiency of air with the combustion of carbon monoxide and 
possibly hydrogen formed in the bed being carried out by the addition of 
additional combustion air above the bed or subsequent to the combustor 14. 
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 18 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 20 or 
overheating of the connecting tube headers and risers 178 180 and 182. 
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 absorbent if sulphur removal 
is required. 
Hot flue gases leave the combustor 14 through the duct 118. These hot gases 
also contain fine dust in suspension and the mixture passes through the 
dust removal system 22. This system 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 
such that 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 6. Ash is removed from separator 22 via pipe 154 
to the ash discharge device 24 which may be a system of lock-hoppers 
and/or rotary valves or similar alternate device. Depressurised ash is 
removed via pipeline 156. Pressurised and cleaned hot gas from separator 
22 passes via duct 120, valve 122 and ducts 124 to the turbine 6 where the 
gases are let down to substantially atmospheric pressure and then pass 
through duct 126 to a waste heat boiler 30 and thence via duct 128 to 
atmosphere. Suitably pressurised feed water is admitted via pipeline 130 
to the boiler 30 which may incorporate a preheater and superheater and the 
resultant produced steam leaves via pipeline 132 for use as described 
below. 
Methane-containing feed gas at a suitable pressure, and which preferably 
has been treated to remove the greater part of undesirable constituents 
such as hydrogen sulphide and all mercaptans, is fed at a controlled rate 
through pipeline 160 to the heat exchange system 38 in which it is heated 
to about 300.degree. C. and thence by pipeline 162 to the zinc oxide 
reactor 40 in which final traces of hydrogen sulphide are removed prior to 
passing via pipeline 164 back to the heat exchange system 38 where a 
controlled amount of steam is added to the gas. The source of the steam is 
described below. The methane/steam mixture is heated to about 500.degree. 
C. in heater 38 before passing via pipeline 166 into the pressurised 
combustor 14. The steam/methane mixture is then distributed by the header 
system 174 and a series of connecting tubes 176 (of which only one is 
shown) to catalyst filled reformer tubes generally indicated at 20. 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 to hydrogen and carbon monoxide and leave the 
bottom of the reformer tubes through connecting stubs 178 (only one is 
shown) to the bottom collection header 180 before passing up the riser 182 
and thence through the shell of the combustor 14 to the heat exchange 
system 38. The riser 182 and the feed line 166 together with the headers 
174 and 180 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 14 using tie rods to support the 
upper tube header 174 the inlet pipe 166 and riser 182. 
The hot reformed gases entering the heat exchange system 38 are cooled and 
in doing so preheat the feed gas as described above and in addition may 
generate steam. Suitably pressurised boiler feed water is provided to the 
heat exchange system 38 via pipeline 168 and steam is generated at a 
pressure suitable for addition to the feed gas as already described. After 
preheating the hot feed gas and steam mixture and generating steam as 
described, there is still sufficient heat in the reformate to preheat 
boiler feed water and generate low pressure steam at about 1.5 to 2.0 Bar 
before being finally cooled by cooling water or air coolers and passing 
via pipeline 184 to the synthesis gas compressor 8. Any additional steam 
for the steam reforming may be supplied, if required, via pipeline 170 and 
in turn this steam may be derived from pipeline 132 and/or 208 described 
later. Low pressure steam leaves the heat exchange system 38 via pipeline 
172. 
Except where CO.sub.2 is present in the feed gas to the reformer, the 
reformate will not normally have the correct stoichiometric 
carbon-to-hydrogen ratio for methanol production. To rectify the 
deficiency, CO.sub.2 is provided from the fluidised bed combustor flue 
gas. In order to extract carbon dioxide from the flue gas a portion of the 
flue gas in pipeline 120 is withdrawn via duct 134 to the heat exchanger 
32 in which the flue gas is cooled to about 200.degree. to 300.degree. C. 
countercurrent to returning scrubbed gas. The cooled gas then passes via 
valve 136 and duct 138 to the scrubbing unit 34 which may contain a final 
cooler and/or direct water wash and packed limestone column to remove 
oxides of sulphur followed by a carbonate solution or equivalent scrubbing 
tower and pressure let down system for the saturated carbonate solution 
and a carbon dioxide separation tower or vessel. The scrubbing system 34 
also includes circulating pumps and solution heat exchangers and coolers, 
make-up solution tanks and feed pumps and any required steam is supplied 
via pipeline 220. Scrubbed gas leaving 34 passes via duct 140 to the heat 
exchanger 32 where it is re-heated to about 800.degree. to 850.degree. C. 
prior to passing via pipeline 14 to join pipeline 124 downstream of valve 
122. The quantity of carbon dioxide produced may be controlled by 
controlling the flow of flue gas to the scrubber 34 by control of valves 
122 and 136 and/or control of the scrubbing conditions, i.e. temperature 
and liquid circulation rate in 34. Carbon dioxide from the scrubbing 
system 34 passes via pipeline 144 to the carbon dioxide compression and 
final purification system 36 in which the carbon dioxide is compressed to 
a suitable pressure and finally purified, if necessary, by suitable means 
such as distillation to remove traces of oxides of sulphur before passing 
via pipeline 146 to pipeline 184 where it is combined with cooled 
synthesis gas to be compressed in compressor 8. Alternatively or 
additionally, suitably treated carbon dioxide at suitable pressure 
available from any alternate source such as an ammonia plant purge may be 
added from pipeline 228, and the addition rate to pipeline 232 and thence 
to pipeline 184 being controlled by valve 230. 
Reformate (with any added CO.sub.2) is compressed by the compressor 8, in 
which intercooling may be employed, and passed via pipeline 186 to be 
combined with re-circulating gas from the methanol synthesis reactor in 
pipeline 198 before passing via pipeline 188 to the recirculation 
compressor 10. Compressed gas from 10 passes via pipeline 190 to the heat 
exchange cooling and separation system 44 in which the gas is heated prior 
to passing to the methanol synthesis reactor 42 via pipeline 192. 
In the reactor, some of the hydrogen and oxides of carbon react to form 
methanol and the reaction gives off heat which boils water in the boiler 
surrounding the reactor tubes. Suitably pressurized feed water is added to 
the boiler via pipeline 206 and steam is withdrawn at up to 40 Bar from 
via pipeline 208. Hot reacted gas passes via pipeline 194 to the heat 
exchange cooling and separation system 44. 
In 44 the hot reacted gases are first heat exchanged with incoming 
synthesis gas and then cooled by water and/or air coolers to separate out 
a crude water-methanol mixture which will also generaly have traces of 
ethers and higher alcohols. The separated liquid is bled off from the 
separator by a control valve to the flash tank, treatment and distillation 
unit 46. In 46 flash gas in the crude methanol is vented via pipeline 213 
as a fuel gas or process gas stream. The crude methanol then may be 
treated with a small controlled amount of caustic added via pipeline 212 
depending on the materials of construction of the distillation plant, and 
the distillation plant then removes light ethers which may be added to the 
vent gas in line 213, water which is removed in pipeline 214, heavy ends 
such as higher alcohols in pipeline 215 and product methanol which is 
removed via pipeline 216. 
Unreacted synthesis gas leaves 44 and passes via pipeline 196 to a branch 
where the bulk of the gas is recycled via pipelines 198 and 188 to the 
recirculation compressor 10. From the branch via pipeline 200 by means of 
valve 202 a controlled amount of purge gas is bled from the system via 
pipeline 204. The purge flow is to remove unreactable components from the 
synthesis loop, one or more inert components such as nitrogen present in 
the feed gas and/or, unreformed methane present after the steam reforming 
reaction and/or hydrogen in excess of that required for the synthesis 
reaction to produce methanol. The purge gas may be used as a fuel gas or 
as a process gas for a separate industry e.g. ammonia manufacture or, if 
the methane content is high, some of the purge may be recycled back and 
added to the feed gas prior to pipeline 160 or the gas may be treated by 
one or more known techniques such as a shift reaction, scrubbing, 
cryogenic separation and adsorption to separate the undesirable inert 
components which then may be vented separately possibly as pure products, 
and the remaining components recycled to the process via pipeline 160 or 
228 as appropriate. 
The methanol plant as a whole is controlled by known and established 
techniques for the control of such plants and gas turbines and as an aid 
to control an alternator 12 is shown producing surplus power via cable 218 
to that required to drive the compressors. This power may be used to drive 
the auxiliary units such as pumps and fans and the power produced by 12 
may be adjusted to compensate for any deterioration in performance of the 
expander 6 between overhauls. 
Alternatively 12 could be an electric motor designed to balance gas 
turbine/compressor assembly 2 in which case power would be supplied via 
cable 218. 
FIG. 2 shows a modification of the arrangement of FIG. 1 suitable for use 
where the steam reforming step is operated at a sufficiently high pressure 
that the reformate in pipeline 184 is provided at substantially the same 
pressure as the unreacted gas in pipeline 198 from the methanol synthesis 
reactor 42. This method of operation permits the replacement of 
compressors 8 and 10 by a single recirculation compressor 8A. 
The reformate in pipeline 184 is combined directly with the unreacted gas 
in pipeline 198 for compression in compressor 8A and subsequent feeding to 
the methanol synthesis reactor 42 via lines 190 and 192. 
Because of the increased pressure of the reforming step, the amount of 
unreacted methane in line 184 is increased, possibly to as much as 10%, 
and a portion of the gas in pipeline 190 is withdrawn through line 302 and 
recycled via valve 304 and pipeline 306 to be mixed with the feed gas in 
160 to the reformer. 
With this arrangement the recirculation compressor 8A circulates synthesis 
gas through the synthesis section 42 and 44 and methane rich purge gas 
through the reformer system 14 and 38 and 40. 
The compressor 8A has an inlet pressure of about 40 to 50 Bar and a 
pressure ratio of about 1.1/1.0 and may be a simple large single wheel 
centrifugal machine. 
Referring to FIG. 3, each 402 is an underground coal seam containing high 
grade coking coal and a significant quantity of adsorbed methane; i.e. 
about 10.0 NM.sup.3 of methane per ton of coal. Two such seams are shown. 
404 is a mine shaft and mine, 406 is a coal washing plant. 408 is a gas 
compression unit 410 is a coal preparation plant, 412 is a shaft sunk for 
coal seam gas drainage. 414 is a gas compression unit and 416 is a 
methanol plant with a coal-fired pressurized fluid bed reformer as shown 
in FIG. 1 or in FIG. 1 as modified by FIG. 2. 
Drill holes 502 are gas drainage holes drilled into the coal seam 402 in 
the operating mine 404 to deplete the seam of gas prior to mining. Gas is 
withdrawn from the coal seam at or slightly above atmospheric pressure and 
is collected in the manifold 504 prior to passing to the gas main 506 
which passes to the above ground compressor system 408 where the gas is 
compressed to about 20-30 Bar or more if the methanol plant 416 is remote 
from the compressor 408. 
Drill holes 510 are gas drainage holes drilled into the coal seams 402 in a 
mine drainage shaft 412 to deplete the coal seams of gas prior to mining 
in order to improve mine safety and/or the rate at which the coal can be 
ultimately mined. Gas is withdrawn from the coal seams at or above 
atmospheric pressure and is collected in manifolds 512 prior to passing to 
the gas main 514 which passes to the above ground compressor 414 where the 
gas is compressed to about 20-30 Bar or more if the methanol plant 416 is 
remote from the compressor 414. 
Coal is withdrawn from the mine 404 by the hoist or conveyor 517 and passes 
via the conveyor 518 to the coal washery 406. In the washery, high grade 
low ash coal is produced and removed via the conveying system 522 as a 
significantly enhanced value product. Washery "middlings" and tailings are 
removed via conveying system 520 and thence via conveying system 526 to 
the coal treatment plant 410. Any excess is conveyed via 524 to dump or 
sale. Coal fed to 410 may contain ash in the range of 30-70% on a dry 
basis and may be supplemented, if necessary, by run-of-mine coal fed to 
410 by conveyor 544 from the mine hoist or coveyor 517. 
In the coal treatment plant 410 the coal is crushed to a suitable size 
range for a fluid bed combustor reformer and may provide the crushed coal 
in dry powder or water slurry form for conveying by the transfer pipeline, 
pneumatic conveyor or alternate device 528 from the treatment plant 410 to 
the methanol plant 416. 
Feed gas to the methanol plant enters via pipeline 516 and 508 from 
compressors 414 and 408 respectively and sufficient drainage shafts 512 
are employed to ensure a reliable and adequate supply of gas. 
The methanol plant 416 operates as described before with reference to FIG. 
1 or 2. Sorbent such as limestone or dolomite as required is supplied via 
conveying system 536; water is supplied via 532 and air is supplied via 
duct 530. Product methanol leaves via pipeline 538, liquid effluent such 
as higher alcohols via pipeline 542 and purge gases which may be used as 
site fuel via pipeline 546, and exhaust gas through vent 540. 
The methanol may be transported by road vehicle or rail using the same mode 
of transport as the coal from the mine or may be pipelined if the coal 
transport should be by slurry pipeline. In the latter case the methanol 
may be either "batched" separately as methanol or as a coal slurry in the 
pipeline. 
EXAMPLE 
Using the plant described and illustrated in FIG. 1 of the drawings, 
methanol is generated from methane gas using the following operating 
conditions. The coal employed to fuel the fluid bed combustor is supplied 
as a slurry in water. Where the 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 employed in the extraction of 
the fuel, e.g. colliery waste water or refinery oily water waste. 
______________________________________ 
Details of steam reformer/fluid bed combustor arrangement 
______________________________________ 
14 
combustors are employed, each being as follows 
(A) Internal diameter of combustor 
5.8 meters 
(B) Depth of fluidised bed 11.0 meters 
(C) Number of reformer tubes per combustor 
500 
(D) Reformer tube length 9.15 meters 
(E) Reformer tube outside diameter 
100 mm 
(F) Temperature of fluid bed 
920 .degree.C. 
Details of Gas turbine assembly 2 
______________________________________ 
(G) Absorbed power of air compressor 4: 
130,000 kW 
(H) Absorbed power of feed gas and recircu- 
lation gas compressors 8,10: 
55,000 kW 
(J) Output of alternator 25,000 kW 
______________________________________ 
__________________________________________________________________________ 
Process flow details 
Flow 
Ref 
Pipeline/Vessel 
Nature rate (Kg/hr) 
Pressure (bar) 
Temperature .degree.C. 
__________________________________________________________________________ 
K line 160 
Methane feed 
39,500 
25 15 
(excluding recycle) 
L *See Below 
Steam to reformer 
140,040 
25 
M Reformer tubes 
reforming steam/ 
183,000 
-- 550-850 
20 methane(including 
recycle methane) 
N Line 184 
reformate 183,000 
20 850 
O Line 228 
CO.sub.2 from 
36,150 
25 15 
external source 
P Line 146 
CO.sub.2 from flue 
nil -- -- 
gas 
Q Line 190/192 
Feed for methanol 
-- 100 -- 
synthesis 
R Line 216 
Methanol product 
104,167 
-- -- 
S Line 104 
Compressed Air 
725,000 
20 
T Line 226 
Coal (as a 60/40 
70,800 
-- -- 
water/coal slurry) 
W Line 124 
HP Flue gas 
902,000 
18 900 
X Line 128 
LP Flue gas 
902,000 
1.01 250 
__________________________________________________________________________ 
*The steam is provided from water supplied through line 168 supplemented 
as required by steam supplied through line 170 from lines 132 and/or 208. 
Total water male up to the plant in lines 130, 206 and 168 is 40,000 
Kg/hr. 
The coal had a calorific value of 5833 Kg Cal/Kg (Y) and was supplied in 
ground form with an average particle size of 0.6 mm (Z). 
By way of comparison, with the same rate of supply of methane to a 
conventional gas-fired steam reformer in which the fuel for the reformer 
is provided in conventional manner from the methane, the rate of methanol 
production would be 64,860 Kg/hr (AA).