Note: Descriptions are shown in the official language in which they were submitted.
:~22~73~
-- 1 --
BACKGROUN~ OF THE INVENTION
2E'IELD OF TE~E INVENTION
3The present invention is directed to an improved
4 process for the production of ammonia synthesis gas, and
specifically to an improved process which utilizes a high
6 hvdrogen content feedstock gas in a single adiabatic re-
7 formin~ stage without requiring a primary refor~ing furnace
8 system as is conventional in present processes.
g DESCRIPTION OF_THF PRIOR ART
Generally the manufacture of ammonia consists of
11 preparing an ammonia synthesis gas from three separate
12 process components: a nitrogen source, usually air; steam;
13 and a hydrogen source which is conventiona~ly either coal,
14 petroleum fractions, or natural gases. For example, in the
preparation of ammonia synthesis gas from a light hydrocarbon
16 feedstock, which may range from natural yas to naphtha, the
17 hydrocarbon feedstock gas is first purified by removing
18 gaseous contaminants, such as sulfur ~which would poison the
19 downstream catalysts) from the feedstock by the catalvtic
hydrogenation of the sulfur compounds to hydrogen sulfide and
21 adsorption of the hydrogen sulfide over a zinc oxide adsorp-
22 tion medium. Subsequent steam reforming of the contaminant-
23 free gas provides the major portion of the hydrogen required
24 ~or ammonia synthesis from the hydrocarbons in the gas.
Reforming is accomplished by a two-stage process in which a
26 mixture of steam and the purified feed gas are first reformed
27 over catalyst in a primary reformer, followed by treatment in
28 a secondary reformer to which air is introduced, in order to
29 provide the required amount of N2 for ammonia synthesis.
30 However, reforming also produces carbon oxides. The carbon
31 monoxide in the refor~ed gas is converted to carbon dioxide
32 and additional hydrogen in a shift conversion step, and the
33 carbon dioxide is removed by subsequent scrubbing. Further
34 treatment of the raw synthesis gas by methanation is con-
ventionally used to remove additional amounts of carbon
373~
-- 2 --
1 dioxide and remaininq carbon monoxide from the hydrogen-rich
2 gas, resulting finally in an ammonia synthesis gas contain-
3 in~ approximately three parts of hydroqen and one part of
4 nitrogen, ~hat is, the 3:1 stoichiometric ratio of hydrogen
to nitrogen in ammonia. The ammonia synthesis gas is then
6 converted to ammonia by passing the gas over a catalytic
7 surface based upon metallic iron (conventionally magnetite)
8 which has been promoted ~ith other metallic oxides, and
g allowing the ammonia to be synthesized according to the
following exothermic reaction:
11 N2 + 3H2 ---3 2N~3
12 The steam reforming of the sulfur-free light hydro-
13 carbon feedstock is conventionally carried out in a two-stage
14 process wherein the first stage, that is primary reforming,
15 produces a partially reformed gas. This partially reformed
16 gas is introduced albng ~ith air into a second stage, that is
17 secondary reforming, to obtain a greater concentration of
18 hydrogen and a lesser concentration of hydrocarbons. The
19 reaction processes occurring in the reforming of the feed-
20 stock gas be~in with the breakdown of hydrocarbons to
21 methane, carbon dio~ide and carbon monoxide:
22 H20 + CnH(2n+2)--~CH4 + CO + C02 + H2
23 and end with the reforming of these products by the desired
24 endothermic methane reforming reaction:
C~4 + H20 ---~ CO + 3H2
26 and by accompanying exothermic reactions:
272CH4 + 7/2 2 ~~~~~ C2 + CO ~ 4H2O
28CO ~ H20------~C02 + H2
292H2 + 2 - 2H2O
30CO + 1/2 2 ~ CO2
31The methane reforming reaction for the production of
32 hydrogen is highly endothermic and, for feedstocks containing
33 less than about 80 vol.% ~2 a large heat transfer is
34 required, which conventionally involves the use of a high
35 capital investment primary reforming furnace, which also
36 consumes a significant amount of energy in the form of fuel.
37 The catalyst for this primary reforming is normally a nickel
38 catalyst supported on alumina.
_ 3_ ~;237~i
1 The suhsequent secondary reforming step takes place in
2 a refractory-lined vessel which also contains a nickel
3 catalyst supported on alumina. In conventional steam
4 reforming processes, air is introduced into this adiabatic
reforming stage to provide the needed nitrogen for the
6 production of ammonia synthesis gas. Oxygen in the air also
7 reacts with the combustion components in the gas stream
8 coming from the primary reforming stage to increase the
g temperature and provide heat for this addi~ional reforming of
10 hYdrOcarbons.
11 U.S. Patent 3,442,613 discloses a two-stage reforming
12 process in which milder primary reforming conditions are
13 employed which results in a larger amount of methane in the
14 primary reEormer effluent. Excess air is then fed to the
15 secondary reformer to permit increased exothermic hydrogen
16 combustion therein,which aids in the reforming of the larger
17 methane volumes fe~ thereto. Thereafter, the excess N2,
18 introduced via the air feed to the secondary reformer, is
19 removed in a cryogenic separation ste~.
U.S. Patent 3,584,998 relates to a one-stage reforming
21 process in which natural gas, excess air and steam are
22 preheated in heat exchange with reformer effluent gas and
23 then reformed, followed by water gas shift and CO2 scrubbinq
24 treatments and then by a cryogenic process in which excess N2
25 is removed from the scrubbed reformer effluent gas.
26 U.S. Patent 3,649,558 also relates to a single stage
27 reformer, in which air is introduced in excess amounts to the
28 secondary~type reformer. Excess N2 is removed in a subsequent
29 cryogenic section.
U.S~ Patents 4,079,017 and ~,162,290 relate to the usa
31 of parallel steam reformers for the primary reforming of the
32 hydrocarbon feed.
33 ~. Chatterjee, "Ammonia From Hydrocarbons--An Improved
34 Processl', Fertiliser News, pp. 19-22 (December 1980) dis-
35 closes another single-stage reforming process in which
36 oxygen-enriched exces~ air is combined with natural gas and
37 steam and reacted in an autothermal reformer, ~ollowed by
38 shift reactions, boiler feed water heat recovery, CO2
_ 4 ~ 3~3~
1 recovery and methanation. Excess ~ethane remains in the
2 reformer ef~luen~ gas and is removed in a downstream cryo-
3 genic section, ~hich also serves to separate excess N2.
SUMMARY OF_THE INVSNTION
6 The Present invention, is broadly directed to an
7 improved process for the production of ammonia synthesis qas,
8 and specifically to an improved process which utilizes a
9 high-hydrogen content feedstoc~ gas in a single adiabatic
10 refo.min~ stage without requiring a primary reforming furnace
11 system as is con~entional in present processes.
12 The improved process of this invention provides an
13 ammonia synthesis gas which, after shift conversion, is
14 characterized by low methane content and does not contain
15 excess nitroqen, and whlch can therefore be passed directly,
16 after conventional treatment for CO2 scrubbing and methan-
17 ation, to an ammonia synthesis reactor for formation of
18 ammonia. The process therefore produces an ammonia syn gas
19 stream without the need for use of the expensive cryogenic
20 purification processes required by the prior art, and at the
21 same time avoids the need to use a conventional Primary
2~ reformer. This results in a large savings in equipment costs
23 and on-going operating expenses.
24 The ~rocess of this invention also permits the use of
25 higher process pressures in the reforming section than have
26 heretofore been possible due to pressure design limitations
27 imposed by current primary reformer tubing metallurgy.
28 BRIEF DESCRIPTION OF THE DRArAINGS
Figure 1 is a schematic illustration of a two-stage
30 primary/secondary reforming process of the prior art.
31 Figures 2 and 3 are a schematic illustration of one
32 embodiment of the process of the present invention.
33 Figure 4 is a schematic illustration of a
34 reactor/exchan~er for use in a second embodiment of the
35 process of this invention.
36 DETAILED DESCRIPTION OF THE INVENTION
37 Referring to the drawings, and specifically the
` ~3~
1 conventional Drimary and secondary reforminq stages which
2 have been hi~hly simplified and illustrated in Figure 1,
3 there is seen a primary reforminq furnace stage generally
4 indicated by the numeral 10, having an upper convection
5 section 8, and a lower primary reforming radiant section 12.
6 The furnace is normally heated by burners mounted in the
7 floor of the radiant section and supplied with fuel gas and
8 combustion air as illustrated. Hot flue gas exitinq the
9 radiant section flows upwardly through convection section 8,
10 past steam superheat exchangers 4 and 6, process air heat
11 exchanger 7, steam generators 9 and 11, steam superheat
12 generator 13, feed gas heat exchangers 15, and 16 and
13 boiler feed water preheater 17, and is discharged through
14 port 3.
16 As discussed previously, conventional two-stage
17 reforming processes require the introduction of four separate
18 2rocess streams to the primary reformer: feed ~as ~for a
19 source of hydrogen), steam, a source of nitrogen gas (which
20 is conventionally process air) and fuel gas. Feed gas is
21 introduced into the process and passes through feed gas heat
22 exchangers 16 and 15, positioned in the primary reforming
23 furnace's convection section 8.
24 This preheats the feed qas to approximately 750F, the
25 require<l temperature for the removal of sulfur over a zinc
26 oxide desulfurization bed 20. Steam is produced by passing
27 boiler feed water through heat exchangers 17, 9, and 11, and
28 by passing the resulting steam for superheating through
29 exchangers 13, 6 and 4, via steam drum 18, as shown, to
30 achieve a steam temperatllre of appros~imately 1200F. This
31 steam is then combined with the desulfurized feed gas and fed
32 to catalyst-filled reformer tubes 14 in radiant section 12.
33 The partially reformed feed gas 28 is then discharged from
34 the primary reformer furnace and fed into secondary adiabatic
35 reformer 30 where it is combined with process air 29 that has
36 been preheated in heat exchanger 7 to about 1200Fo The
37 oxyqen in this preheated process air reacts with combustibles
38 (H2, CO, and CH4) in the partially reformed feed gas, and
39 releases additional heat. t~pon entering the secondary
- 6 ~ 3~
1 reformer's catalyst bed the gas mixture undergoes additional
2 reforming with a decrease in temperature due to the more
3 predominant endotherrnic reactionO The raw synthesis gas is
4 discharged from the secondary reformer and under~oes addi-
tional processing in conventional steps: carbon dioxide is
6 formed from carbon monoxide in shift converter unit 40;
7 carbon dioxide is removed in process unit S0; carbon monoxide
8 and additional carbon dioxide is removed by methanation in
g process unit 60; the synthesis gas is compressed in com-
pressor 70 to that required for ammonia synthesis in the
11 ammonia synthesis system 80; vaporous ammonia undergoes
12 further compression in refriqeration compressor 90; and is
13 withdrawn from the process as ammonia ~roduct. Excess heat in
14 the ammonia synthesis section is removed by means of boiler
feed water heat exchanger 82.
16 Referring to ~igure 2, wherein one embodiment of the
17 process of this invention is illustrated, the selected
18 high-hydrogen content feedstock is passed via conduit 108 to
19 ~irst heat exchanger l10 wherein it is heated, generally to a
temperature of from about 300 to 350F, by indirect heat
21 exchange with a low temperature shi~t converter effluent gas,
22 to be described in more detail below, which is passed to
23 exchanger 110 via conduit 201.
24 The gas feedstocks which can be treated in accordance
with the process of this invention for sinqle-stage reforming
26 are gases containing high concentrations of hydrogen, i.e.,
27 hydrogen concentrations greater than about 40 vol.%, and
28 preferably greater than about 50 vol.~ H2, and most typically
29 from about 50 to 80 vol.~ H2, in addition to lower hydro-
carbons, small concentrations of carbon monoxide, and trace
31 gases, such as CO2, H2S, COS, N2 and argon. The lower
32 hydrocarbons present in the gas feed generally comprise
33 members selected from the group consisting of saturated
34 aliphatic hydrocarbons having from 1 to 4 carbon atoms, and
unsaturated aliphatic hydrocarbons having from 2 to 4 carbon
36 atom~; are principally methane but also inclusive of
37 ethylene, ethane and the like; and are generally present in a
38 concentration of from about 15 to 30 vol.~. The carbon
_ 7 ~ 36
l monoxide concentration in the gas feed is not critical and
2 will generally range from about 0 to 25 vol.~. Among these
3 potential high-hydrogen content feedstocks are coke gas or
4 refinery gases, such as are discussed in ~.S. Patent
3,649,558, in addition to coal pyrolysis gas, and feedstocks
6 such as those available from an intermediate BTU gas (I~G)
7 streams resulting from the gasification of coal or lignite
8 using conventional gasification processes.
g The heated feed gas is withdrawn from exchanger 110 via
conduit 112 and may be admixed with sufficient steam (which
ll can be introduced to conduit 112 via conduit 111) to supply
12 the water of reaction required for a subsequent COS hydroly-
3 sis reaction (which can be effected in COS hydrolysis reactor
14 114), if COS is present in the feed. The quantity of steam
which is thus introduced can vary widely and will generally
16 comprise from about 2 to 4 vol.~, based on the total feed
7 gas in conduit 112 withdrawn from exchanger 110. The COS
18 reaction in reactor 114 can be effected by any conventional
l9 means, using conventional hydrolysis catalysts such as
activated alumina. In this reactor, COS contained in the feed
21 gas is converted into hydrogen sulfide gas at conventional
22 hydrolysis conditions, which typically range from abo~t 300
23 to 350 F and from about 300 to 600 p5ig. If the feedstock
24 gas does not contain COS, steam injection line 111 and COS
25 hydrolysis reactor 114 can be eliminated from the system if
26 desired.
27 The resulting gas mixture is withdrawn via conduit
28 116 and is introduced into second heat exchanger 118 wherein
29 the gas is further heated, in this case by indirect heat
30 exchange with a high temperature shift converter gas effluent
31 (to be described in more detail below~, which is introduced
32 thereto via conduit 164. Thereafter, the further-heated gas
33 which may contain the hydrogen sulfide and which will
34 generally have a temperature of from about 700 to 750F, is
35 withdrawn via conduit 117 and introduced into sulfur removal
3~ zone 120, wherein the hydrogen sulfide impurities are removed
37 from the gas stream by conventional technology, such as hy
38 use of a zinc o~ide adsorption bed. The gas, now essentially
- 8 ~ 73~
1 free of sulfur impurities (e.g., containing less than about
2 0.2 ppm by weight of sulfur compounds, calculated as elemen-
3 tal sulfur), is withdrawn via conduit 122 and admixed with
4 steam, which can be accomplished by injected steam into
conduit 122 via conduit 124. Again, the quantity of steam
6 introduced at this point can vary, and will generallly range
7 from about 2.5 to 4.0 moles per mole of carbon in the
8 desulfurized gas feed. The function of the steam introduced
~ at this point in the process is to provide the water of
reaction necessary for the subsequent reforming reactions.
11 The steam/desulfurized gas mixture is then further heated
12 (generally to a temperature of from about 800 to 1,000F) in
13 third heat exchanger 126 by indirect heat exchanger with a
14 portion of the reactor/exchanger effluent gas (to be des-
cribed in more detail below) which is introduced to exchanger
16 126 via conduit 138. The thus-heated steam/desulfurized gas
17 mixture 128 is introduced into the tube side of tubular heat
18 exchanger 130 wherein the feed gas is at least partiaïly
19 reformed by contacting the feed gas, under reforming condi-
tions, in tubes 121 of reactor/exchanger 130 with a conven-
21 tional reforming catalyst. Any conventional primary
22 reforming catalyst can be employed, such as nickel, nickel
23 oxide, chromia, molybdenum, mixtures thereof and the like,
24 with nickel-on-calcium aluminate or nickel-on-alumina being
preferred. The temperature within tubes 121 will generally
26 range from about 900 to 1500F and the pressure will gener-
27 ally range from about 300 to 1000 psig, and the total gas
28 hourly space velocity in tubes 121 will generally range from
29 about 5000 to 15,000 v/v/hr., with a range of from 8000 to
10,000 v/v/hr. being preferred.
31 As a result of the reforming reactions occurring in
32 tubular exchanger 130, substantially all of the hydrocarbon
33 component~s of the feed gas (other than methane) are converted
34 to CH4, CO, CO~ and H2; a portion of the original methane
35 components are likewise converted to CO, CO2 and E~2; and the
36 temperature of the gas mixture will be qenerally increased to
37 about 1300 to 1450F. The partially reformed gas will
~ ~23t7~
g
1 generally have a residual methane level of from about 5 to
2 20 vol.% CH4, on a dry basis.
3 Process air obtained from any convenient source i5
4 passed via conduit 197 to fourth heat exchanger 196 wherein
it is heated (generally to a temperature of from about 700 to
6 800F) by indirect heat exchange with a ~ortion of the high
7 temperature shift converter effluent gas which is passed
~ thereto via conduit 166. The thus-heated process air is
9 withdrawn via conduit 195 and passed to yet another heat
exchanger 194 for further heating (generally to a temperature
11 of from about 900 to 1000F) by indirect heat e~change with a
12 portion of the cooled reformer effluent gas, which is passed
3 thereto via conduit 135 from the shell side of exchanger 130.
14 The thus-heated process air is then introduced via conduit
141 into secondary reEormer 140, together with the partially
16 reformed gas mixture which is introduced via conduit 136.
17 The ~uantity of air introduced via conduit 141 is
18 adjusted using conventional control means (not shown), to
19 provide an air:feed ratio sufficient to yield about a 3:1
hydrogen:nitrogen ratio in the ammonia synthesis gasl that
21 is, to pro~ide a H2:N2 ratio of from about 2.6:1 to 3.2:1,
22 and preferably from about 2.8:1 to 3.1:1.
23 Secondary reformer 140 comprises an adiabatic reformer
24 of conventional design and can be provided with suitable
internal burners to be used during start-up of the process in
26 order to bring the temperature within the reformer to within
27 the range of from about 1400 to 1600F, after which further
28 heating can be accomplished via the heat released in the
29 exothermic reaction of oxyqen therein with feedstock. The
30 amount and type of catalyst in reformer 140 is also con-
31 ventional, with Ni catalysts supported on alumina being
32 typical. The secondary reformer will qenerally employ a
33 temperature of from about 1600 to 1900F, a pressure of from
34 about 300 to 1000 psig, and a total gas hourly space velocity
35 o~ from about 7000 to 10,000 v/v/hr.
36 The reformer efluent gas (generally having a tem-
37 perature of from about 1600 to 1800F and a residual C~4
38 level of from about 0.2 to 0.6 vol.% CH4, on a dry basis) is
- 10 ~ 3~
withdrawn from secondary reformer 140 via conduit 134 and is
2 passed to the shell side of reactor/exchanger 130 for
3 indirect heat exchange with, and heating of, the
4 steam/~esulfurized feed gas mixture pa<;sed to exchanger 130
via conduit 128, as described above. ~rtle effluent gas
5 withdrawn via conduit 132 is then divided into two portions.
7 A first portion is passed via conduit 138 to third heat
8 exchanger 126 for indirect heat exchange with, and heating
9 of, the steam/desulfurized gas mixture as described above.
10 The second portion is passed via conduit 135 to fifth
ll exchanger 194 to provide the final stage of heating of the
12 process air in conduit 195 prior to its injection into
13 secondary reformer 140. The partially cooled reformer
14 effluent gas i9 withdrawn from third exchanger 126 via
15 conduit 139 and passed to steam superheater 142 and first
16 steam generator 158, superheater 142 receiving steam via
17 conduit 144 ~generally at a temperature of from about ;90 to
18 600 F and about 1500 psig) from steam drum lS0 and producinq
19 superheatefl steam which is withdrawn via conduit 143
20 ~generally at a temperature of about 800 to 900F and about
21 1500 psig), and generator 158 in turn generating steam 159
22 from water stream 157 which is fed thereto from steam drum
23 150. From generator lS8, the cooled reformer effluent is
24 passed via conduit 155 to high temperature shift converter
25 160, wherein carbon monoxide in the reformer effluent gas is
26 converted over conventional catalysts and using conventional
27 methods and equipment to carbon dioxide and additional
28 hydrogen. Partially cooled effluent gas is also withdrawn
29 via conduit 137 from fifth exchanger 194 and is recombined
30 with the remaining reformer effluent in conduit 155.
31 Generally, a temperature of from about 700 to 900F and
32 a pressure of from about 300 to lO00 psig will be employed in
33 converter 160, and the catalyst will generally comprise a
34 supported, chromium-promoted iron catalyst. Thereafter, gas
35 exiting the high temperature shift converter is withdrawn via
36 conduit 162 and is itself split into two portionsO A first
37 portion is passed via conduit 164 to second heat exchanger
38 118 for heating of the qas feed to desulfurization zone 120,
3'7~
1 as described above. The partially cooled effluent gas is
2 ~hen withdrawn via conduit 165 from exchanger 118 and passed
3 to low temperature shift converter 200, preferably af~er
4 treatment in guard bed 190.
The second portion of the gaseous effluent from ~he
6 hiyh temperature shift converter 160 i9 passed via conduit
7 169 to a second steam generator 152 in which steam 151 is
8 produced from water feed 153 and is returned to drum 150 from
g which water 153 was received. The partially cooled high
temperature shift effluent from generator ~52 is then itself
11 split into two portions: a first part is passed via conduit
12 166 to fourth heat exchan~er 196 to provide the first staqe
13 Of heating of the process air, introduced thereto via conduit
1~ 197, as described above. The further cooled effluent gas is
then withdrawn via conduit 167 and passed to conduit 1fiS for
16 combined feed to the low temperature shift converter 200, or
17 preferably first to guard bed 190.
18 The second part of shift effluent from exchanger 152 is
19 passed via conduit 168 to boiler feedwater exchanger 170 in
which boiler feedwater, introduced thereto via conduit 180~
21 is heated and from which the further cooled shift effluent is
22 withdrawn (via conduit 174) and combined with stream 165 for
23 feed to low temperature shift converter 200, or preferably
~4 first to guard bed 190. If desired, feedwater 180 can be
first heated by exchange ~ith low temperature shift effluent
26 201 in a separate exchan~er (not shown) prior to introduction
27 to exchanger 170.
28 Guard bed 190, which is optional, is preferably
29 employed to treat gas stream 165 upstream of low temperature
30 shift converter 200 to remove halide and sulfur impurities
31 and thereby protect any halide- and sulfur-sensitive catalyst
32 in low temperature shift converter 200. Tha operation of
33 guard bed 190 is conventional and is generally conducted at
34 temperatures and pressures within the ranges used in low
35 temperature shift converter 200 as described oelow, and the
36 solids employed in guard bed 190 for such halide and
37 S-impurities removal generally comprise the same catalyst as
38 is used in low temperature shift converter 200.
1 In shift converter 200, a low temperature shift
2 conversion reaction is effected over conventional catalyst
3 usinq conventional methods and equipment to form additional
4 quantities f ~2 and CO2. Generally, a temperature of fro~
5 about 400 to 500F and a pressure of from about 300 to 1000
6 psig ~ill be employed in converter 200, and the catalyst will
7 generally comprise a mixture of zinc o~ide and copper. The
8 effluent gas from low temperature shif~ converter 200 is
g passed via conduit 201 to first heat exchanger 110, as
10 described above, for heating of the feed gas introduced
11 thereto via conduit 108. The cooled, low temperature shift
12 converter effluent gas, now depleted of its heat values, is
13 then withdrawn via conduit 106 and (referring now to Figure
14 3) can be passed to CO2-removal zone 250, in which any
15 conventional process (e.g., solvent absorption of CO2 gas)
16 can be used to remove CO2 via conduit 210. The resulting
17 CO2-free gas is fed by conduit 220 to conventional methana-
18 tor zone 300 for removal of additional CO and C2 and is then
withdrawn (via conduit 320), compressed in compressor 350 and
20 passed as direct feed via conduit 370 to ammonia synthesis
21 zone 400,-wherein NH3 is formed from the H2/N2 synthe-
22 sis gas 370 (i.e., 3:1 ~2:N2 molar ratio) using conven-
23 tional techniques ~i.e., over Fe-catalyst at 700 to 950F).
24 Waste gases are withdrawn via conduit 410 and product N~3 is
25 recovered via conduit 420~
26 The o~eration of CO2 removal zone 250, methanation zone
27 300, compressor 350 and N~3 synthesis zone 400 is conven-
28 tional and need not be more completely described for a full
29 understanding of the process of this invention. The precise
30 operating parameters and equipment of each such process step,
31therefore, will be readily apparent to one having ordinary
32skill in the art, and each step can include the usual
33internal recycle streams and stages found useful in the prior
34art. Thus, CO2-removal zone 250 can include conventional
3sCO2-absor2tion and CO2-desorption stages wherein the
36CO2-laden gas 106 is contacted with a liquid contalning
37 either a solven~ for, or a dissolved compound (e.q.~ R2CO3)
38 readily reactive with, the CO2; the CO2-free gases (generally
~3~3~;
-- 13 --
l containing less than about 0.15 vol.~ CO2) are withdrawn; and
2 the solvent is treated to desorb the CO2 gases 210 for
3 recycle of solvent to the absorber. Exemplary of suitable
4 conventional CO2 removal systems are those discussed in
5 ~irk-Othmer, Encyclopedia of Chemical Technology, 3d Ed.,
6 Vol. 2, pp. 492-494 ~1978). Similarly, methanator 300 will
7 generally employ a temperature in the ran~e of about 570 to
8 940F, a pressure from about 300 ~o 1000 psig, and a sup-
g ported Ni catalyst (e.g., Ni on alumina) to convert any
o remainin~ CO and C02 in ~as stream 220 to rnethane, thereby
ll producing an effluent gas 320 containing less than about 10
12 vppm (i.e., parts per million by volume) of total CO and CO2
13 and H2 and N2 in a H2:N2 molar ratio o~ from about 2.6:1 to
14 3.1:1. Compression in zone 350 can take ?lace in several
lS stages, as desired, to bring the methanator effluent to
16 synthesis reactor pressure, which generally ranges from about
17 2000 to 5000 psig. Finally, ammonia synthesis zone 400 can
18 include conventional dryer units wherein trace water i5
19 removed from the syn ga9 as required and conventional purge
20 r~covery units wherein a portion or all of the gas effluent
21 from the ammonia synthesis reactor is treated to recover and
22 recycle H2 to the reactor and to rernove inerts such as C~4
23 and Ar therefrom.
24 The improved process of this invention produces a syn
5 gas 370, having a H2:N2 ~nolar ratio of about 3:1, that is a
26 H2:N~ molar ratio of from about 2.6:1 to 3.2:1, and which has
27 a residual methane concentration (dry basis) of less than
28 about 2 vol.~, and more typically less than about 1 vol.~ and
29 which is therefore suitable for direct feed to an ammonia
30 synthesis reactor zone 400, that is a syn gas 320 which is
31 not subjected to a cryogenic purification followinq methana-
32 tor 300 to remove excess methane prior to the ammonia
33 synthesis reaction. The elimination of the cryogenic section
34 required by the prior art and the avoidance of the use of a
35 primary reformer furnace by the improved process of this
36 invention results in a very large savings in terms of
37 equipment investment and operating expense.
3~ In accordance with another embodiment of the process of
39 this invention, illustrated in ~igure 4, steam/desulfurized
.
~.~23~3~
~ 14 -
l gas mixture 12aa is passed to the shell side of reactor/
2 exchanger 130a and the selected reforming catalyst is
3 housed in the shell side of apparatus 130a. The steam/
4 desulfurized gas mixture is at least partially reformed
5 over the catalyst whilq beinq heated by means of reformer
6 effluent gas 134a, which in this embodiment is passed to the
7 tube side 121a of reactor/exchanger 130a. The partially
8 re~ormed gas 136a and the partiall~ cooled reformer e-ffluent
9 gas 132a are withdrawn and further treated as described above
lO for streams 132 and 136, respectively, in the e~bodirnent
ll illustrated in Pigure 2. The conditions of operation of
12 apparatus 130a in ~igure 4 correspond to those discus~sed
13 above for apparatus 130 in the embodiment illustrated in
14 Figure 2.
To further illustrate the process, a feed gas con-
16 taining 52 vol.~ H2, 20 vol.~ CO, 28 vol.% CH4 and 30 vppm
17 COS, having a temperature of about llO~F and a pressure of
18 about 350 psig, is preheated to about 315F in tubular feed
~ gas preheat exchanger 110, and is then introduced into COS
20 hydrolysis drum 114, after addition of 2 vol.~ steam ~750F,
21 600 psiq), based on total feed gas 112, twithdrawn from
22 exchanger 110), in which the carbonyl sulfide is converted to
23 hydrogen sulfide over a bed of an alumina hydrolysis catalyst
24 (at 315F, 350 psig, 2000 v/v/hr. qas hourly space velocity).
25 The temperature of the resulting feed qas is increased to
26 about 750F, the temperature required for ~urther desulfur-
27 ization, by passing the hydrogen--sulfide-containing gas
28 through tubul~r heat exchanger 118, followed by adsorption of
Z9 the hydrogen sulfide over a zinc oxide adsorption bed 120.
To the sulfur-free feed gas discharged from zinc oxide
31 adsorption bed 120 is then added 3.0 moles of steam (750F,
32 600 psig) per mole of carbon in the feed gas, and this
33 combined stream passes through tubular heat exchanger 126 in
34 order to increase the gas temperature to about 900F.
8, An additional increase in the temperature of this
36 combined stream (to about 1400F) is obtained by subsequent
37 preheat treatment in the tube side 121 o~ reactor/exchanger
38 130 in indirect heat exchange with reformer effluent gas
36
- 15 -
l enterin~ the shell side of exchanger 130 and in contact with
2 nickel on alu~ina reforming catalyst in tubes 121 (at 1375F,
3 325 psi~, 9000 v/v/hr.) so that partial steam reforming of
4 hydrocarbons takes place in exchanger 130, further con-
tributing to the preheat of this feed stream to be charged to
6 reformer 140.
7 ~rocess air is adjusted using conventional control
8 means (not shown) to provide an air:feed ratio sufficient to
9 yield about a 3:1 ~2:N2 ratio in the final product ammonia
synthesis gas 320. The adjusted process air enters the
ll process by first being pressurized (compressor not shown) to
12 about 50 psi above the pressure used in reformer 140. The
13 pressurized process air is then preheated in tubular heat
14 exchanger~ 196 and 194 to a reforming inlet temperature of
about 1000~F.
16 The heated process air feed and the steam-feed gas
17 streams are then introduced into adiabatic reformer 140
18 wherein reforming of the feed gas takes place over Ni on
l9 alumina reforming catalyst (at a space velocity of about 7000
v/v/hr.)~
21 The reformer effluent gas from reformer 140 (1730F,
22 320 psig) is discharged, as discussed above, into the shell
23 side of exchan~er 130, and from exchanger 130 (1130F, 315
24 psig) is passed to feed preheat exchanger 126 and air preheat
25 exchanger 194. The partially cooled reformer effluent from
26 feed gas preheat exchanger 126 is passed to steam superheater
27 142 and first steam generator 158 for generation of 1500 psiq
28 steam, and the resulting cooled reformer effluent withdrawn
29 from generator 158 is combined with the remaining cooled
reformer effluent 137 and passed via conduit 155 (700F, 310
31 psig) to high temperature shift converter 160. In converter
32 160, a water gas shift reaction is effected over an
iron-based catalyst (720F inlet temperature, 310 psig) to
34 form ~2 and CO2 from CO contained in the reformer effluent
35 gas. Gases e.xiting converter 160 (850F, 300 psig) are then
36 passed to heat exchangers 118, 196 and 170 and steam
37 generator 152 (qenerating steam at 1500 psig) as described
38 above, combined in conduit 165 (420F, 300 psig) and fed to
- 16 - ~2~3~
l low temperature converter guard bed 190 containing ZnO/Cu
2 solids to absorb any Cl and S values which may be present,
3 followed by treat~ent in low temperature shift converter 200,
4 containing conventional ZnO/Cu shift conYersion catalyst.
After passing through heat exchanger 110, the product gases
6 (at 450F, 290 psi~) are withdrawn via conduit 106, and
7 treated for CO2 removal in CO2 removal zone 250, methanated
8 in methanator 300, compressed and then fed to ammonia
9 synthesis zone 400, as described above. The ammonia synthesis
l0 gas withdrawn from methanator 300 contains H2:N2 in about a
ll 3:1 molar ratio and contains less than about 1.0 vol.~ CH4
12 (dry basis), and less than about 13 vppm CO and CO2.
13 A feedstock 108 suitable for the single stage reforming
14 process of the present invention is, for example, a steam
15 cracker tail gas having a hydrogen content of about 70~, with
16 the remainder being methane, or an IBG stream from a coal
17 gasification process having major component composition of
18 hydrogen (60~), carbon monoxide ~20%), and methane(16~).
l9 Still another feedstock suitable for the .single stage
20 reforming process of the present invention, and that used as
21 the feedstock material for obtaining the data in Table 1 has
22 the composition oE hydrogen (52.13~), carbon monoxide
23 ~19.39%), methane (27.03~), carbon dioxide ~0~), nitrogen
24 (0~), and mixed 2-carbon hydrocarbons (1.45%).
A comparison of the parameters is tabulated in the
26 following Table 1 for (1) a conventional process as shown in
27 Figure 1, except that the primary refor~er is omitted,
28 conducted at the maximum practical preheat temperature for
29 feedstock and air; and (2) single stage reforming using the
30 feed/effluent reactor/exchanger 130, as depicted in Figure
31 2.
- 17~ 3~73~i
~0 z 0~ O ~ o o
o ~ i~l o o o ~
O O ~ P~ ~ ~ ~ ~
Z ~ Z~ ~ o
, ~
~ S ~ a ~ ~ O ~ ~ ~ c . .
3 _ V 3
C
o ' ~ O ~
X o
S e _ ~ 4 ~ 1 e
V 0 1.. ~ O ~ Q tJ O = ~D 2 0
~ ~ ~ O ae oh O a~ a~ ~ cn o 0
V ~ O ~1 ~ O 0 3
. ~ :.'
- 18 - ~373~
l Table 1 clearly indicates that a feed/steam preheat to
~ 1375F in the reactor/exchanger and accompanying partial
3 reforming of hydrocarbons therein results in a decrease in a
4 final product syn gas methane content from an unacceptably
high 5.86 vol.~ to 0.34 vol.%, thereby indicating the greater
6 reforming efficiency obtained utilizing the reactor/exchanger
7 130.
8 When the amount of feed gas required to produce a
9 constant amount of hydrogen in the ammonia syn gas is
calculated, the efficiency of the total process is easily
ll seen. For example, as shown in Table 1, 0.65 lb-mols/hr. of
12 f eed gas are required to produce 1.0 lb-mols/hr. of hydrogen
13 in the product syn gas in the reactor/exchanger process
14 design depicted in Figure 2, whereas 0.8a lb-mols/hr. of feed
15 gas are required to produce the same amount of hydrogen under
16 the conventional process desLgn depicted in Figure l when the
17 primary reformer is omitted.
18 The advantages of the process design according to the
l9 present invention are therefore seen as a reduction in
capital expenditure associated with the construction of a
21 primary reforming furnace, as otherwise required to produce
22 an acceptable concentration of methane in product synthesis
23 gas, an economic savings in utility costs expended in the
24 heating of said primary reforming furnace, a desirable
25 decrease in the methane slippage of the ammonia syn gas
26 produced, and a greater efficiency in utilizing the feed gas
27 requirements to yield the ammonia syn gas produced.
28 Thus, while I have illustrated and described the
29 preferred embodiment of my invention, and have described my
30 invention and the manner and process of making and using it
31 in such full, clear, concise and exact terms as to enable any
32 person skilled in the art to which it pertains to make and
33 use the same, one skilled in the art can easily ascertain the
34 essential characteristics of this invention and without
35 depar~ing from the spirit and scope thereof can make various
36 changes and/or modifications to the invention Lor adapting it
37 to various usages and conditions. Accordingly, such changes
38 and/or modifications are properly intended to be within the
39 full range of equivalents of the following claims.
