Note: Descriptions are shown in the official language in which they were submitted.
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Process of Producing Methanol from Natural Gas
This invention relates to a process of producing methanol
from natural gas, the natural gas is catalytically converted
with steam and molecular oxygen, and a raw synthesis gas is
withdrawn from the catalytic conversion, which rw synthesis g~
contains H2, CO and C02, a gas rich in H2 and recycle gas are admixedto
the raw synthesis gas and a synthesis gas mixture is pro-
duced, which is passed through at least one methanol synthe-
sis reactor which contains a copper catalyst and is operated
at a pressure in the range from 10 to 150 bar and at tempera-
tures in the range from 180 to 350°C, from the methanol syn-
thesis reactor a gas mixture containing methanol vapor is
withdrawn, which is cooled to such an extent that methanol
condenses out, a cooled residual gas is separately withdrawn
from the methanol and divided into the recycle gas and a sec-
ond partial stream (purge gas), from the purge gas, which has
a HZ content of at least 30 vol-% (calculated dry), a gas
rich in H2 is separated in a separating means and admixed to
3D the raw synthesis gas.
Such process is known from EP-B-0 233 076. The natural gas or
a partial stream of the natural gas is first of all passed
through a steam reforming, and the cracking gas thus produced
together with the residual natural gas is then charged into
an autothermally operated reactor which contains a catalyst.
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Steam reforming is effected as usual in a tubular furnace,
where a nickel catalyst is provided in a plurality of tubes,
which are disposed in a combustion chamber and are heated
from the outside by hot combustion gas. The steam reforming
in accordance with the known process serves to increase the
hydrogen content in the succeeding autothermal reactor and
also to provide for an increased hydrogen content in the raw
synthesis gas produced.
The object underlying the invention is to modify the known
process such that the process can be performed with low in-
vestment costs, and at the same time a favorable consumption
of process means is achieved. It should also be possible to
work without steam reforming. In accordance with the inven-
tion this is accomplished in the above-mentioned process in
that
(a) the catalytic conversion of natural gas with steam and
oxygen is performed only in the autothermally operated
reactor, which comprises at least one burner fed with the
natural gas and the oxygen, and below the burner a bed of
granular catalyst;
(b) the raw synthesis gas is withdrawn from the autothermally
operated reactor at a temperature in the range from 800
to 1200°C, is cooled, steam is condensed out, and conden-
sate produced is separated;
(c) the raw, cooled synthesis gas has a stoichiometric number
S of 1.3 to 1.9 prior to admixing the gas rich in H2,
where S is calculated from the molar concentrations of
H2, CO and Co2 in accordance with
S = (H2 - C02) . (CO + C02);
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(d) the purge gas consists of 30 to 80 vol-% H2,
7 to 35 vol-%, and preferably
at least 10 vol-% COZ and
3 to 25 vol-% CO,
(e) the hourly quantity of purge gas is 8 to 25 % vol-%, and
preferably 10 to 15 vol-% (calculated dry) of the hourly
quantity of cooled, raw synthesis gas;
(f) the hourly quantity of CO2 supplied to the separating
means in the purge gas is 10 to 50 %, and preferably 15
to 35 % of the hourly quantity of C02 in the raw synthe-
sis gas; and
(g) the hourly quantity of Co supplied to the separating
means in the purge gas is 2 to 10 % of the hourly quan-
tity of CO in the raw synthesis gas;
(h) the gas rich in H2, which was withdrawn from the separat-
ing means and admixed to the raw synthesis gas, has a H2
content of at least 80 vol-%, and by admixing H2 a stoi-
chiometric number S of 1.95 to 2.2 is achieved for the
synthesis gas.
In the process in accordance with the invention it is ensured
that the purge gas represents a relatively large gas stream.
This is achieved in that the methanol synthesis reactor is
operated such that apart from the methanol vapor a large re-
sidual gas stream can continuously be withdrawn. It is delib-
erately omitted to achieve an optimally high consumption of
H2, CO and C02 in the synthesis reactor. In particular the
consumption of C02 is reduced, so that the consumption of H2
is also reduced considerably. In the process in accordance
with the invention it is ensured that the purge gas quantity
contains at least the amount of hydrogen required for the
necessary increase of the stoichiometric number of the raw
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synthesis gas. In the methanol synthesis reactor the neces-
sary amount of catalyst decreases, so that the methanol syn-
thesis reactor can be designed smaller. It is also possible
to operate the reactor at a relatively low pressure or with a
low recycle ratio. Apart from these advantages considerably
reducing the costs of the plant it should be noted that the
process is also conducted without CO conversion and without
C02 removal in the raw synthesis gas. Since steam reforming
is omitted in accordance with the invention, it is definitely
possible to employ a higher pressure in the autothermally op-
erated reactor. Since the pressure in the methanol synthesis
reactor can at the same time be relatively low, it is thus
possiblelto largely adjust the pressures in the gas produc-
tion and in the synthesis.
The autothermally operated reactor, which operates without
indirect heating of the catalyst, is known and described for
instance in Ullmann's Encyclopedia of Industrial Chemistry,
5th edition, Vol. A 12, pp. 202 to 204. The natural gas is
supplied to the reactor preferably preheated. Usually, tech-
nically pure oxygen is supplied to the burner of the reactor,
in order to keep the content of inert gas in the raw synthe-
sis gas as low as possible. The supply of steam usually lies
in the range from 0.5 to 3.0 mol, based on the molar carbon
content of the natural gas. The autothermally operated reac-
tor can have a one-part or a multistage design. Instead of
natural gas, the autothermally operated reactor can also be
charged with other charges, e.g. with liquefied gas or refin-
ery gas.
For the methanol synthesis reactor known constructions are
considered, in particular the water-cooled tubular reactor or
the adiabatically operated fixed-bed reactor.
The separating means for producing the gas rich in H2 is
likewise designed in a manner known per se, e.g, as pressure-
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swing adsorption plant or low-temperature gas separation
plant.
Embodiments of the process will now be explained with refer-
ence to the drawing. The drawing represents a flow diagram of
the process.
The autothermally operated reactor 1 comprises a bed 2 of
granular nickel catalyst. Above the bed, a burner 3 is dis-
posed, to which desulfurized natural gas is supplied through
line 4, gas rich in 02 is supplied through line 5, and steam
is supplied through line 6. As gas rich in o2 there is nor-
mally used technically pure oxygen. The heat required for the
conversion in the reactor 1 is only applied through partial
oxidation. The temperatures at the outlet 8 of the reactor 1
lie in the range from 800 to 1200°C, and usually are at least
900°C.
Hot, raw synthesis gas leaves the reactor 1 and flows through
a cooling 9, which may have a multistage design, which is not
represented in detail in the drawing. Upon cooling, an aque-
ous condensate is produced, which is withdrawn via line 10.
The cooled, raw synthesis gas, which flows through line 11,
usually has temperatures in the range from 20 to 80°C. This
gas is compressed by a compressor 12. Through line 14, gas
rich in H2 is supplied, which comes from a separating means
15. The H2 content of the gas in line 14 is at least 80 vol-
%, and preferably at least 90 vol-%.
The stoichiometric number S of the gas in line 11 lies in the
range from 1.3 to 1.9, and mostly is not more than 1.8. By
admixing the gas rich in H2 of line 14, the stoichiometric
number of the synthesis gas in line 16 is increased to 1.95
to 2.2, as this is necessary for methanol synthesis. S is
calculated in the known manner from the molar concentrations
of
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H2, CO and C02 in accordance with S = (H2 - C02) . (CO + C02)~
To the gas in line 16, recycle gas is admixed via line 18,
and the synthesis gas mixture thus formed is first of all
supplied to a heat exchanger 20 via line 19, before it enters
the methanol synthesis reactor 22 through line 21.
In the drawing, the reactor 22 is represented as known tubu-
lar reactor, where the copper catalyst is provided in a plu-
rality of tubes 23. For the indirect cooling of the interior
of the tubes 23, cooling water is supplied to the reactor
through line 24, and heated and partially evaporated cooling
medium is withdrawn via line 25. In the tubes 23, the tem-
peratures lie in the range from 180 to 350°C, and mostly in
the range from 200 to 300°C. The pressure in the tubes is 10
to 150 bar, and mostly lies in the range from 20 to 100 bar.
From the reactor 22 a gas mixture containing methanol vapor
is withdrawn via line 26, which gas mixture is first of all
cooled in the heat exchanger 20 and is then supplied to an
indirect cooler 28 via line 27. Via line 29, the mixture is
supplied to the condenser 30, from which a condensate con-
taining methanol and water is withdrawn via line 31. The fur-
ther distillative treatment of this condensate, which is
known per se, need not be explained in detail.
From the condenser 30, cooled residual gas is withdrawn
through line 32 and is divided into two gas streams. The
first stream is recirculated as recycle gas through line 18.
The second stream, here referred to as purge gas, is supplied
to the separating means 15 via line 33, which separating
means operates for instance according to the principle of
pressure-change adsorption. In the separating means 15 an ex-
haust gas containing methane, CO and C02 is produced in addi-
tion to the gas rich in H2 of line 14, which exhaust gas is
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withdrawn via line 34 and can be used as combustion gas or be
supplied to some other use.
For admixing the gases of lines 14 and 18 to the synthesis
gas, a compressor will be required, which was, however, omit-
ted in the drawing for simplification. The reactor 1 can be
operated at a pressure above 40 bar, where the compressor 12
can be omitted.
Examples:
The following examples, which are calculated in part, are di-
rected to the production of 1000 tons methanol per day. The
nickel catalyst in the reactor 1 consists of 15 wt-% nickel
on a carrier containing A12o3. The copper catalyst for the
methanol synthesis contains 60 wt-o copper, 30 wt-o zinc ox-
ide and 10 wt-o aluminium oxide and operates at a pressure of
80 bar. The separating means 15 is a pressure-swing adsorp-
tion plant, for which a H2 yield of 85% is assumed.
Example 1:
There is used a plant corresponding to the drawing; the natu-
ral gas in line 4 has been preheated to 500°C, just as the
technically pure oxygen in line 5 and the steam in line 6.
Per hour, 1472 kmol natural gas, 1946 kmol steam, and 850
kmol 02 are supplied to the burner 3. In addition to methane,
the natural gas contains 1.5 mol-o C2H6 and 0.4 mol-% C3Hg +
C4Hlo. Further details are shown in the following Table I,
where in different lines the components of the respective
mixture are indicated in mol-% and kmol/h:
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Table I
Line I 8 I 14 I 16 I 33 I 34
mol-% kmol/hmol-% mol-% mol-% kmol/hkmol/h
COZ 6.1 383 - 8 12 60 60
CO 17 1070 - 22.4 4.9 24 24
H2 47.3 2982 100 69 73.4 365 56
CH4 0.8 50 - 1 8.9 44 44
H20 28.8 1821 - _ _ - _
CH30H - - - - 0.3 2 2
Total amo~lnt kmol/h6310 310 4800 498 188
Tem erature C 1000 35 35 35 35
Pressure bar 35 83 82 7g 5
The stoichiometric number S of the raw synthesis gas in line
8 is 1.79, and for the gas in line 16 S is 2Ø
Example 2_
The pressure in the reactor 1 is now 50 bar, so that the com-
pressor 12 can be omitted. There is used the same natural gas
as in Example 1, the temperature in lines 4, 5 and 6 is each
600°C. Per hour, 1498 kmol natural gas, 1982 kmol steam and
800 kmol technically pure oxygen are supplied to the burner
3. Further data are listed in Table II:
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TnlW n, T T
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Line 8 14 16 33 34
mol-% kmol/h mol-% mol-% mol-% kmol/h kmol/h
COZ 5.9 372 -- 7.8 10.7 50 50
CO 16.7 1057 -- 22.2 6.3 30 30
HZ 47.3 2993 100 67.8 61.6 289 54
CH4 1.7 103 -- 2.2 20.8 95 95
H20 28.4 1794 -- -- 0.1 -- --
CH30H __ __ __ __ 0.5 2 2
Total amount 6233 235 4764 470 234
kmol/h
Tem erature 1000 35 35 35 35
C
Pressure bar 50 53 52 49 5
S 1.83 2
1st Comparative Exam le (calculated):
For the production of 1000 tons methanol per day there is
used the natural gas of Example 1, preheated to 500°C,
whereof 360 kmol/h are delivered through a commercial tubular
furnace for steam reforming to a commonly used nickel cata-
lyst. Thus, a H2-containing primary gas with about 40 vol-
H2 is produced, to which 1030 kmol/h natural gas are admixed.
The mixture is charged into an autothermally operated reac-
tor, as it is solely used in the inventive process in accor-
dance with the drawing. By adding 02 and steam, a raw synthe-
sis gas is produced in a total amount of 5900 kmol/h, with a
temperature of 1000°C and a pressure of 35 bar, where S =
1.95. Through line 14, 66 kmol/h hydrogen are admixed to this
gas, and S is thus increased to 2Ø The amount of purge gas
(line 33) is only 147 kmol/h, which is charged into a pres-
sure-change adsorption plant. Through line 34, 81 kmol/h ex-
haust gas are discharged, which is composed of (in kmol/h)
8.2 C02, 4.1 CO, 21.1 H2, 0.5 methanol, 2.4 N2, and 44.7 CH4.
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It should be noted that in particular the quantity and compo-
sition of the purge gas are considerably different from Exam-
ples 1 and 2.
2nd Comparative Example (calculated):
There is used the natural gas of Example 1 and the process in
accordance with the drawing. From the raw synthesis gas of
line 11, 200 of the Co2 are removed, so that the stoichiomet-
ric number is increased to 1.93. The different lines contain
the following gas quantities and gas compositions:
Line 11 14 16 33 34
mol-% kmol/hmol-% mol-% mol-% kmol/h kmol/h
COz 7.1 307 -- 6.9 6.1 9 9
CO 24.2 1050 -- 23.7 3 5 5
HZ 67.5 2924 100 68.2 65.8 98 3
CH4 1.1 49 -- 1.1 23.1 35 35
CH30H + Nz -- -- -- -- 2 3 3
Total amount 4330 95 4425 15 55
kmol/h
Tem erature 35 35 35 35 35
C
Pressure bar 34 83 82 7g 5
S 1.93 2
Here as well, the composition and quantity of the purge gas
in line 33 is considerably different from Examples 1 and 2.
