Note: Descriptions are shown in the official language in which they were submitted.
l~lQ664
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T~IIS INV~TI02~ ~ELAI~S to a process for producing methanol by the
catalytic reaction of one or more carbon oxides with hydrogen.
~he reaction of carbon oxides with hydrogen to give methanol is
exothermic.
C0 + 2E2 ) CH30H ~lI = -21685 kg cal/mol
C2 + 3~2 ~ CH30E ~ H20 ~H = -11830 kg cal/mol
and therefore in prinoiple a methanol synthesis process should be capable of
providing a quantity of usable heat. In modern methanol synthesis processes
using a copper-containing catalyst, however, the highest temperature obtained
by the reacting mixture of carbon oxides and hydrogen is usually under 300 C
and rarely above 270C. Consequently it is not practicable by passing such
mixture through a waste-heat boiler to raise steam at a pressure greater than
about 50 ata. Steam at such a relatively low pressure can, of course, be made
use of; and, indeed? processes have been proposed in which steam is raised in
a special reaotor in which the catalyst is disposed in the tubes of a boiler or
; boiler tubes are disposed between layers of catalyst. The disadvantages enter
in, howerer, that turbines in which such steam can be let down for power
recovery are thermodynamically limited in efficiency as compared with higher-
pressure turbines. 'rurbines o~ the condensation type may be used but these
are higher in oapital cost than the pass-out turbines employed when higher-
pressure steam i9 generated, as in many ammonia plants. Moreover the special
catalytic reactors are complicated and expensive.
A methanol production plant normally includes, in addition to the synthesis
section, a synthesis gas generation section in which a carbonaceous feedstock
is con~rerted to carbon oxides and hydrogen by a high temperature reaction with
steam and/or oxygen. We have realised that by integrating in a special way
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the heat recovery in the synthesis gas generation section a highly efficient
over-al:L heat recovery can be obtained with less incidence of the above-
mentioned disadvantages.
According to the first aspect of the invention there is provided a
methanol production process which comprises(a) generating methanol synthesis gas in one or more stages in at least one
of which there is produced a gas stream at over 400 C;
(b) generating steam at a pressure of at least 50 ata, by heat exchange with
such strea~ or streams;
(c) bringing synthesis gas to synthesis pressure by means of a compressor
powered from an engine in which such steam is let down;
(d) synthesising methanol over a catalyst at an outlet temperature of under 300 C;
(e) transferring heat evolved in the synthesis to water maintained under a
pressure too high to permit boiling to take place;
(f) passing the resulting hot water to stage (b) as feed for the steam generation;
and
(g) recovering methanol from the cooled gas rom stage (e).
Methanol synthesis gas generation usually involves the reaction of a
carbonaceous feedstock, such as natural gas, refinery off-gas, gaseous hydrocarbons,
non-vaporisable hydrocarbons, coal or coke, with steam and possibly also carbon
dioxide or oxygen. ~he reaction of such materials takes place typically at
over 700 C and may be as high as 1100 C for a catalytio process, still hlgher
for a non-catalytic process, in order to effect sufficiently complete reaction
to crude synthesis gas containing carbon oxides and hydrogen. If the feedstock
is one of the first 4 mentioned the reaction ls most often carried out without
oxygen over a catalyst in tubes externally heated in a furnace ("steam reforming")
but can be carried out in an insulated vessel if oxygen is also fed ("partial
oxidation"). If the feedstock is one of the last 4, the reaction is usually
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carried out in the presence of oxygen without a catalyst. Depending on the
hydrogen-to-carbon-ratio of the carbonaceous feedstock and on the extent
which oxygen is used, synthesis gas generation may involve a C0-shit and
C02-removal stage to bring the hydrogen to carbon oxides ratio to the level
required for methanol synthesis. ~he crude synthesis gas is cooled and freed
from its content of unreacted steam before passing it to the synthesis section.
Synthesis gas generation may alternatively begin with the shift reaction
of carbon monoxide with steam to give carbon dioxide and hydrogen (outlet
temperature over 400C) and C02-removal, if carbon monoxide is available as
a starting material.
The pressure in the synthesis gas generation section is typically up to
100 ata and thus the gas usually has to be compressed before feeding it to the
methanol synthesis.
The streams by heat exchange with which steam is generated in stage (b)
include the orude synthesis gas stream and the flue gas of the furnace if a
steam reforming process is used. The steam pressure i8 preferably in the
range 80-120 ata, as a result of which it is practicable to let it down in an
engine of the pass-out type and to use the exhaust æteam as the feed for the
synthesis gas generation section. The engine may drive the synthesis gas
compressor directly or may drive an electric generator powering the compressor.
In favourable conditions enough steam can be generated to provide, directly or
indirectly~ the mechanical power required in other parts of the process, such
as the synthesis gas circulator (if a recycle process is used) and various
feed-pumps and fans. It is within the invention, howev'er, to raise some of
the steam in a fired boiler or by burning fuel in the flue-gas duct of a
reformer furnace, and to use some of the waste-heat steam in condensing engines
or in engines exhausting at less than synthesis gas generation pressure~ for
example into the re-boiler of a methanol distillation.
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Af-ter the waste-heat boiler and the economiser associated with it for
the steam generation, the temperature of the streams of crude synthesis gas
or reformer furnace flue gas is suitably in the range 200-300 C and preferably
more than 225 C. This can be higher than is typical of methanol processes
previously proposed because the water fed to the economiser has been heated
(for example to 200-260C) by heat evolved in the synthesis instead of merely
being warmed (for example to 140-180C) by further heat exchange with crude
synthesis gas. As a result, other strearns can be heated by the crude
synthesis gas, in particular the hydrocarbon feed to the synthesis gas
generation sec~ion and/or purge gas from the synthesis, especially if ~ is to
be let-down in an engine according to the second aspect of the invention
described below. A further result of water-heating by heat evolved in the
synthesis is that the temperature differences across the boiler and economiser
can be smaller than were previously used, and thus they can be smaller units.
Thu~ the capital cost of the added heat exchangers is in part repaid by the
lo~rer cost of the boiler and economiser.
After heating the other streams the crude synthesis gas or reformer
furna~e flue gas is typically at 140-180 C and can warm the boiler feed water
to be heated by heat evolved in the synthesis and can raise low pressure
steam before being cooled below the dew-point of the steam contained in it.
The methanol synthesis at under 300 C can be at any convenient pressure.
Recently developed processes at 50 ata or 100 ata are very suitable as part
of the process of the invention, but lower and higher pressures, for example
in the range 30-400 ata can be used. The catalyst for such processes usually
contains copper and also zinc oxide and one or more further oxides, such as
chromium oxide, as described for example in our ~K specification 1 010 871 or
oxides from Groups II-IV of the Periodic Table (especially of aluminium) as
described for example in our ~K Specification 1 159 035, or possibly of
manganese or vanadium.
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A variety of general types of methanol synthesis process have been
proposed, differing in the methods adopted for handling the heat evolved
in the synthesis reaction. Any one or more of these can be used excepting,
of course, those designed to use directly all the relatively low pressure
("intermediate pressure") steam generated by heat exchange with the reacting
gas or reacted gas in the synthesis. ~hus synthesis may be over a catalyst
in tubes surrounded by a coolant or in the space around tubes containing
coolant~ The coolant may be for example pressurised water or a mixture of
diphenyl and dipher~l ether; the pressurised water can be used as feed for
the high pressure steam generation or, like the mixture, heat-exchanged in
liquid form with boiler feed water to be fed to the high pressure steam
generation. Alternatively the coolant water may be allowed to boil and the
res~lting intermediate pressure steam condensed in heat exchange with the
water to be fed to the high pressure steam generation. In another process
-the catalyst bed can be in several p æts with heat-abstractior. by coolant
between the parts. In a third process the catalyst temperature can be con-
trolled by heat exchange with cool feed gas passing through tubes in the
catalyst bed or through the space surrounding catalyst-filled tubes. For
the first two of such processes reactors not much simpler than previously
proposed steam-raising processea are required, however, and it may therefore
be preferred to use the third or, better still, a process in which the
temperature is controlled by injecting cool synthesis gas ("quench gas")
into the hot reacting synthesis gas. Quench gas can be injected into mixing
chambers between successive parts of a catalyst bed or~successive reactor
vessels. A very convenient system involves a single body of catalyst in
which are disposed catalyst-free perforated hollow bars each having a sp æ ger
for introducing the quench gas, the bars being l æge enough in cross section
for their interiors to constitute mixing zones and close enough together
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B 27123 A
or to the catalyst bed walls to cause a substantial proportion of reaction
mixture to pass through their interiors, as described in our UK specification
1 105 614. The temperature of quench gas can be below 50 C, but therm~l
efficiency is better if it is at between 50 and 150C, as will be discussed
below.
The volume space velocity of the flow of gas through the catalyst bed
is typically in the range 5000-50000 hour 1 and is preferably fixed at a
level such that the gas leaves the catalyst bed when the quantity of metnanol
formed has been sufficient to raise the gas temperature to the design level,
which is under 300 C and most preferably under 270 C. The methanol content
of the reacted gas is for example 2-5% for a process at 50 ata and
proportionately more at higher pressures. Consequently unreacted carbon
oxides and hydrogen are left over after methanol has been recovered and are
preferably passed again over a methanol synthesis catalyst, for example, by
recirculatiDn to the inlet of the catalyst and mixing with fresh synthesis gas~
~he above space velocity range refers to the mixture in such a process.
In a preferred way of transferring to the feed water for high pressure
steam generation the heat evolved in the synthesis, reacted gas leaving the
catalyst is passed through two parallel heat exchanges, the first of which
heats synthesis gas to synthesis inlet temperature, wbich i~ preferably 20-
40& lower than the outlet temperature of the catalyst bed. Tbe seoond heats
water to a temperature preferably in the range 200-260 C under a pressure
too high to permit boiling to take place or heats a coolant (such as aescribea
above) f~om which heat is to be transferred to such wa~er. ~he reacted gas
becomes cooled initially to 150-190 C in these exchangers. Preferably it i9
then (suitably after re-uniting the two streams) heat-exchanged with cold
synthesis gas from the generation section or methanol recovery or both. This
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affords a useful secondary heat recovery and decreases
the capacity required of the first heat exchanger
After secondary heat recovery the gas is passed to a
cooler and separator for recovery of methanol
In the alternative way of transferring heat to
the feed water, by raising steam in the reactor and
condensing it in heat exchange with the feed water,
the reacted gas leaving the reactor can be cooled to
50-150C in a single heat exchange with cold synthesis
gas and then passed to the cooler and separator,
Unreacted gas from the separator is preferably
recirculated but, if the fresh gas has a hydrogen to
carbon oxides ratio different from stoichiometric
and/or contains non-reactive gases such as nitrogen,
methane or argon, it is necessary to purge a part of
it in order to prevent the concentration of such gases
from building up too much in the gas passing over the
catalyst. Since the purge gas is at only slightly
under synthesis pressure, a useful energy recovery
results from letting it down in an expansion engine
Since the purge gas is at the low temperature of
methanol separation, it is capable of absorbing low-
grade heat from other process streams in the plant and
thus the energy recovery from purge gas is yet more
valuable, After letting-down, the purge gas can be used
as a fuel or source of hydrogen for purposes such as
feedstock desulphurisation,
Such let-down of purge gas, especially after low-
grade heat absorption, constitutes a second aspect of
the invention, applicable also in methanol production
processes outside the scope of the statement of the
first aspect of the invention.
10664
Although the first aspect of the invention resides
essentially in transferring the heat evolved in methanol
synthesis to water without boiling it, it is within the
invention to conduct part of the synthesis so as to
raise steam directly
The first aspect of the invention is applicable
to a methanol production process operated in conjunction
with ammonia synthesis by making a nitrogen-containing
crude synthesis gas and using the methanol synthesis
purge gas as feed for the ammonia synthesis section.
The drawings show two flowsheets of processes
according to the invention: -
Figure 1 shows heat recovery from reacted
synthesis gas directly as boiler feed water; and
Figure 2 shows generation of intermediate pressure
steam in the synthesis reactor, followed by heating
boiler feed water by condensation of such steam.
Both figures show po~er recovery by letting down
synthesis purge gas through a turbine.
Synthesis qas qeneration section (common to both flowsheets).
Reformer 10 includes catalyst-filled tubes 11
suspended in a refractory lined box heated by burning
natural gas (burners not shown) and having a flue gas
duct 12 in which are disposed heat exchangers 14 A-E.
Exchangers A-D will be referred to in relation to the
streams to be heated in them. Exchanger E is a combustion
air preheater for the natural gas burners. The feed to
reformer 10 is a mixture of steam and desulphurised natural
gas which has been preheated in exchanger 14A.
(Desulphurisation is by known means and is not shown).
664
Over the catalyst reaction occurs to give crude synthesis
gas containing carbon oxides and hydrogen and excess
steam. This gas is cooled in waste-heat boiler 16 and
then in economiser 20, both of which with heat exchanger
14C, serve high-pressure steam drum 18, The gas is cooled
further in parallel exchangers 22 and 2~; in 22 it transfers
heat to methanol synthesis purge gas and in 24 to natural
gas to be mixed with steam. From these exchangers the
gas passes to boiler feed water heater 26, cooler 28
(which may include a low-pressure boiler) and water-
separator 30.
Methanol synthesis section as shown in Fiqure 1,
After separation of water at 30 the gas is compressed
centrifugally by compressor 32 and mixed therein at an
intermediate pressure level with recirculated gas from
methanol separation, ~he mixed gas is divided at 33
into 2 streams, one of which is heated in exchangers
34 and 36 and fed to the main inlet 38 of synthesis reactor
40; and the other of which is fed without heating to
the quench inlets 42 of reactor 40. (If desired, the
gas stream can be divided between exchangers 34 and 36
and warmed gas fed to quench inlets 42). Quench inlets
42 suitably lead to spargers each disposed within a hollow
bar having perforations small enough to prevent catalyst
particles from entering but large enough to cause gas
to pass from the catalyst bed into the bars so that it
mixes with quench gas, Reacted gas heated by the
exothermic synthesis reaction leaves reactor 40 and is
divided into two streams, one of which passes through
the hot side of exchanger 36 in which it heats incoming
3 11~)664
synthesis gas and the other of which passes through boiler
feed water heater 44 in which it heats further the water
that has been warmed in heater 26 and is to be passed
via economiser 20 to high-pressure steam drum 18,
The streams leaving exchanger 36 and heater 44 are re-
united and passed through the hot side of exchanger 34
in which cold synthesis gas is warmed. The gas is cooled
to methanol condensation temperature in cooler 46.
Methanol is recovered in separator 48. The unreacted
gas leaving separator 48 is divided at 50 into a
recirculation stream to be passed to the intermediate
pressure section of compressor 32 and a purge stream
to be treated for energy recovery by heating in
exchangers 22 and 14D and letting down in turbine 52.
The power requirements of compressor 32 and the
various other machines employed in carrying out the pro-
cess are supplied by purge-gas let-down turbine 52,
steam turbine 54 (high pressure pass-out) and steam
turbine 56 (low pressure pass-out or condensing).
Direct drives may be used or some or all of the turbines
may generate electricity to be used in electric motor
drives or, in favourable conditions to be exported.
Process example based on flowsheet of figure 1.
The heat recoveries in the process are illustrated
by the stream temperatures (in degrees C) shown on the
flowsheet, These relate to a process using 1600 kg
mol~hour of natural gas as process feed and 91 metric
tons/hour of steam at the inlet of reformer tube 11 and
producing 41.665 metric tons/hour of methanol. The
pressure at the exit of reformer tube 11 is 20 ata.
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664
and compression is to 102,3 ata at the inlet of reactor
40. The compositions and flow-rates of the gases in
the synthesis section are as sho~n in Table l, -
The improvement in thermal efficiency resulting
from the first aspect of the invention is based on the
heat exchanged between reacted synthesis gas and boiler
feed water in item 44, such that warm water (155C)
from exchanger 26 is heated to 237OC before being fed
to the economisers 20 of the high pressure steam system.
Since heating to 237OC is effected in the synthesis
section, the sensible heat of the crude synthesis gas
leaving economiser 20 is available for an intermediate
level of heat recovery by exchange with purge gas at 22
and feed natural gas at 24. The improvement in thermal
efficiency resulting from the s0cond aspect of the
invention is based on the let-down of purge gas from
a pressure of 94 ata in turbine 52, after being the
recipient of waste heat from synthesis gas in exchanger 22
and flue gas in exchanger 14D.
Methanol sYnthesis section as shown in fiqure 2
After separation of water at 30 the gas is compressed
centrifugally at 32 and mixed in the compressor at an
intermediate pressure level with recirculated gas ~rom
methanol separation. The mixed gas is heated in heat
exchanger 58 to synthesis inlet temperature and fed to ~ -
the inlet of synthesis reactor 60 in which it passes
over methanol synthesis catalyst contained in tubes 61,
which are surrounded by water. As the synthesis proceeds,
heat is evolved and is absorbed by the water, which passes
up into drum 64, where it boils, while liquid water is
fed into the reactor shell at 62 to replace it.
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Reacted gas leaves reactor 60, passes through the hot
side of heat exchanger 58 in which it gives up heat to
cold gas from compressor 32, and is then cooled cooled
to methanol condensation temperature in cooler 46.
Methanol is recovered in separator 48. The unreacted
gas leaving separator 48 is divided at 50 into the
recirculation stream to be passed to the intermediate
pressure section of compressor 32 and a purge stream
to be treated for energy recovery by heating in
exchangers 22 and 14D and letting down in turbine 52,
Steam generated in drum 64 is divided at 68 into two
streams, One of these is passed to boiler feed water
heater 70 in which condensation takes place in heat
exchange with water that has been warmed in heater 26
and is to be passed via economiser 20 to high pressure
steam drum 18. The other stream is exported. Part of
the water warmed in heater 26 is fed with the condensed
steam to drum 64 at 72.
The power requirements of compressor 32 and the
various other machines employed in carrying out the
process are supplied in the same way as for the process
of figure l.
Process example based on flowsheet of fiqure 2
The heat recoveries in the process are illustrated
by the stream temperatures (in degrees C) shown on the
flowsheet. Apart from the slightly lower temperature
of the gas leaving item 26, the temperatures are the
same as in figure 1, for the synthesis gas generation
section. The compositions and flow rates of the process
gases are the same as in the process of figure l and
are set out in Table 1.
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The improvement in thermal efficiency resulting
in the process of ~igure 2 from the first aspect to ~he
invention is based partly on the heat recovered as ste~m
in reactor 60 and transferred to boiler feed water in
item 70, such that warm water (155C) from exchanger 26
is heated to 237C before being fed to the econo~isers
20 of the high pressure steam system, As in the process
of figure 1, the sensible heat of the crude synthesis
gas leaving economiser 20 is available for an intenmeaiate
level of heat recovery by exchange with purge gas at 22
a~d feed natural gas at 24, The over-all thermal
ef~iciency is rather better than that obtained using
the process of figure 1 since the reacted gas entering
the cooler is at 99C instead of 120C, so that less
heat is discharged to atmosphere in cooler 46, The
fuel consumption is, however, the same as in the process
of figure 1, the greater e~ficiency being exploited in
the form of exported intermediate pressure steam, as
shown in Table 2,
TABLE 1
.
Composition % V/v ¦ Flow rate
Gas ICO ¦Co2 H2 C~ ~C UeO~ ~2 ¦ Rm3/hour
. _ . .
sis gas 5.,g 6.4 73,1 3,9 0,0 _ 0,6 149800
Reactor feed 4,8 2.7 79.8. 10.7 0,03 0,2 1.7 710040
Reactor Outlet 1.7 1,6 76.5 11,8 1.4 5.1 1.9 646724
Purge 1.8 1.7 81,6 12,6 O.O O,3 2.0 45424
14
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The improvement in thermal efficiency due to the
first aspect can be illustrated by considering the sources
of the heat required to produce the hi~h-pr~ssu P stea~
(145 metric tons/hour, 100 ata 530QC) from water at 110C,
as shown in Table 2. If the second aspect of the invention
is used, as in the flow-sheet, a further 4.0 x 106 kg cal/
hour are recovered.
TABLE 2
_ . ~ ........................................ __
. .. Quantity of heat, 106 kg cal/hour
. . . ..... ~
Source of heat... .... Pre.v.ious process Invention process
. _ . _ _ . ____
Cooling reformer 62.~5 55.53
gas from 850C
Synthesis gas at .
44 (directly or at _ 12 68
70 (via steam). .
Reformer gas
low-grade heat _ 6.92 .
Total recovered 62.45 75.13
. ~ . .. . . .
Flue gas or 40.935 28.255 .
extra fuel ... ... . .. .. _ ___
Total xequired 103.385 103.385
. . __ .. . ,
Export steam, 50 ata _ 8.050
. tfigure 2 only) .___ _ . . - .
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Both aspects of the invention are applicable to processes
. in which methanol synthesis is combined with further reactions,
such as the formation of dimethyl ether, hydrocarbons or
: oxygenated hydrocarbons.
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