AMMONIA SYNTHESIS

SYNTHESE DE L'AMMONIAC

English Abstract

AMMONIA SYNTHESIS
ABSTRACT OF THE DISCLOSURE

An improved ammonia synthesis train utilizing hydrocarbon
starting materials which are converted to hydrogen under
superatmospheric pressure in a series of steps which include
conversion of the hydrocarbons to hydrogen-carbon monoxide
mixtures, shift conversion of the carbon monoxide to hydrogen
and CO2, and removal of the CO2. In such a system, reduction
in process gas losses is achieved through an improved arrange-
ment and integration of the shift conversion and CO2 removal
steps. The carbon monoxide is first shifted in a series of
shifts including a final low temperature shift; the bulk of
the CO2 is removed by an essentially isothermal hot potassium
carbonate CO2 removal system; the small residual amount of
carbon monoxide is shifted to very low levels in a final low
temperature shift; and the small amount of residual CO2 is
then removed to very low levels in a second non-isothermal
scrubbing system employing an aqueous solution of a chemical
absorbent. Methanation, compression and ammonia synthesis
follows. Losses of process gas in the methanation step and in
the purge gas from the ammonia loop are sharply reduced with
concomitant high thermal efficiency.

Claims

THE EMBODIMENTS OF THE INVENTION IN WHICH AN EXCLUSIVE PROPERTY OR PRIVILEGE
IS CLAIMED ARE DEFINED AS FOLLOWS:
1. In a process for the synthesis of ammonia from a hydrocarbon
starting material involving the steps of generating under superatmospheric
pressure a mixture comprising carbon monoxide and hydrogen from said
hydrocarbon, converting the carbon monoxide in said mixture to hydrogen and
CO2 by reaction with water through catalytic shift conversion, removing
said CO2 from said mixture by absorption in a liquid scrubbing medium which
is regenerated by steam stripping, converting residual CO and CO2 to methane
by reaction with hydrogen over a methanation catalyst, compressing said
hydrogen together with nitrogen in about the molar ratio required for ammonia
synthesis, passing said compressed mixture over an ammonia synthesis
catalyst, recycling unconverted synthesis gas to the ammonia synthesis
catalyst, and purging sufficient gas from the recycle stream to maintain a
desired concentration of methane and other inerts in the stream flowing to
the ammonia synthesis catalyst, a method for minimizing hydrogen losses in
the methanation step and hydrogen and ammonia losses in the purge step
while at the same time minimizing the steam required by such CO2 removal
and CO shift conversion steps which comprise the steps of:
a. Converting from 90 to 99% of the carbon monoxide in said carbon
monoxide-hydrogen mixture to hydrogen and CO2 by reaction of said
carbon monoxide with water in at least two successive catalytic
shift conversion zones, the last of which operates in the
temperature range of from 350 to 550°F;
b. Removing carbon dioxide from the mixture produced in Step (a)
in a first CO2 scrubbing zone to provide a gas mixture containing
not more than about 2% and not less than about 0.1% CO2, the CO2
removal in said first zone being carried out by contacting said
mixture with a scrubbing solution comprising an aqueous solution
of potassium carbonate, wherein at least the major portion of said
scrubbing is carried out at a scrubbing solution temperature in the
38

vicinity of the atmospheric boiling temperature of said solution,
and wherein the regeneration of said scrubbing solution is carried
out at a reduced pressure by steam stripping thereof, whereby the
bulk of said CO2 is removed with a minimum consumption of stripping
steam;
c. Subjecting the gas mixture produced in Step (b) to catalytic
shift conversion at a temperature in the range of from 350° to 550°F
to convert residual carbon monoxide to hydrogen and carbon dioxide
to produce a gas stream containing not more than 0.1% residual
carbon monoxide;
d. Removing CO2 from the gas mixture produced in Step (c) in a
second CO2 scrubbing zone to provide a mixture containing not more
than about 200 ppm of residual CO2, the CO2 removal in said second
zone being carried out by contacting said gas mixture with a
scrubbing solution comprising an aqueous solution of an alkaline
chemical absorbent wherein said scrubbing is carried out at an
absorption temperature of from 90° to 140°F permitting the
reduction of residual CO2 in said mixture at least to said level
of 200 ppm and wherein the regeneration of said scrubbing solution
is carried out by steam stripping thereof;
e. Converting the residual carbon monoxide and carbon dioxide in
the gas mixture from Step (d) to methane in a catalytic methanation
zone, and thereafter introducing said gas mixture into an ammonia
synthesis loop including a purge for preventing the build-up of
inert gases in said loop.
2. A process in accordance with claim 1 in which the steam for
the steam stripping of said scrubbing solutions in Steps (b) and (d) is
supplied by transferring heat present in said gas mixture to said scrubbing
solutions.
3. A process in accordance with claim 1 in which the process gas
stream is heated and saturated with water prior to the final low temperature


39

shift conversion in Step (c) by direct contact with a circulating stream of
water which is heated in turn by direct contact between the circulating water
stream and the hot process gas.
4. A process in accordance with claim 1 wherein the hot process
gas leaving Step (a) is cooled prior to Step (b) by direct contact with a
circulating stream of water and wherein the heated water thereby generated
is utilized to heat and saturate the process gas stream leaving Step (b)
by directly contacting such process gas stream with said heated water prior
to introducing said mixture into the final low temperature shift conversion
zone in Step (c).
5. A process in accordance with claim 1 in which from 95% to
98% of the carbon monoxide in the carbon monoxide-hydrogen mixture is con-
verted to hydrogen and CO2 in Step (a).
6. A process in accordance with claim 1 in which the removal of
carbon dioxide in Step (b) is carried out to provide a gas mixture containing
not more than about 1% and not less than about 0.2% CO2.
7. A process in accordance with claim 1 in which the CO2 removal
in Step (d) is carried out to provide a mixture containing not more than
about 100 ppm of residual CO2.
8. A process in accordance with claim 1 in which the temperature
of the last shift conversion in Step (a), and the temperature of the shift
conversion in Step (e) is in the range of from 400° to 500°F.
9. A process in accordance with claim 1 in which the absorption
temperature in Step (b) is in the range of from 220° to 250°F.
10. A process in accordance with claim 1 in which the absorption
temperature in Step (d) is in the range of from 100° to 130°F.
11. A process in accordance with claim 1 in which the scrubbing
solution employed in Step (d) is an aqueous solution of an alkanolamine.
12. In a process for the synthesis of ammonia from a hydrocarbon
starting material involving the steps of generating under superatmospheric
pressure a mixture comprising carbon monoxide and hydrogen from said


hydrocarbon, converting the carbon monoxide in said mixture to hydrogen and
CO2 by reaction with water through catalytic shift conversion, removing said
CO2 from said mixture by absorption in a liquid scrubbing medium which is
regenerated by steam stripping, converting residual CO and CO2 to methane
by reaction with hydrogen over a methanation catalyst, compressing said
hydrogen together with nitrogen in about the molar ratio required for ammonia
synthesis, passing said compressed mixture over an ammonia synthesis catalyst,
recycling unconverted synthesis gas to the ammonia synthesis catalyst, and
purging sufficient gas from the recycle stream to maintain a desired concen-
tration of methane and other inerts in the stream flowing to the ammonia
synthesis catalyst, a method for minimizing hydrogen losses in the methanation
step and hydrogen and ammonia losses in the purge step while at the same
time minimizing the steam required by such CO2 removal and CO shift conversion
steps which comprise the steps of:
a. Converting from 95 to 98% of the carbon monoxide in said carbon
monoxide-hydrogen mixture to hydrogen and CO2 by reaction of said
carbon monoxide with water in at least two successive catalytic
shift conversion zones, the last of which operates in the
temperature range of from 400° to 500°F;
b. Removing carbon dioxide from the mixture produced in Step (a)
in a first CO2 scrubbing zone to provide a gas mixture containing
not more than about 1% and not less than about 0.2% CO2, the CO2
removal in said first zone being carried out by contacting said
mixture with a scrubbing solution comprising an aqueous solution of
potassium carbonate, wherein at least the major portion of said
scrubbing is carried out at a scrubbing solution temperature in the
vicinity of the atmospheric boiling temperature of said solution,
and wherein the regeneration of said scrubbing solution is carried
out at a reduced pressure by steam stripping thereof, whereby
the bulk of said CO2 is removed with a minimum consumption of
stripping steam;

41


c. Subjecting the gas mixture produced in Step (b) to catalytic
shift conversion at a temperature in the range of from 400° to 500°F
to convert residual carbon monoxide to hydrogen and carbon dioxide
to produce a gas stream containing not more than 0.05% residual
carbon monoxide;
d. Removing CO2 from the gas mixture produced in Step (c) in a
second CO2 scrubbing zone to provide a mixture containing not more
than about 100 ppm of residual CO2, the CO2 removal in said second
zone being carried out by contacting said gas mixture with a
scrubbing solution comprising an aqueous solution of an alkaline
chemical absorbent wherein said scrubbing is carried out at an
absorption temperature of from 90° to 140°F permitting the
reduction of residual CO2 in said mixture at least to said level
of 100 ppm and wherein the regeneration of said scrubbing solution
is carried out by steam stripping thereof;
e. Converting the residual carbon monoxide and carbon dioxide
in the gas mixture from Step (d) to methane in a catalytic
methanation zone, and thereafter introducing said gas mixture
into an ammonia synthesis loop including a purge for preventing
the build-up of inert gases in said loop.


42

Description

. 1(~3~59`~
BACKGROUND OF THE INVENTION AND PRIOR ART
This inventio~ relates to improvement in ammonia synthesis tralns.
In current practice almost all synthetic ammonia is produced with the
use of hydrocarbon feedstocks to furnish the source of the hydrogen required i-~
in the cataly~ic ~ynthesis step, the necessary nitrogen being derived from air.
To minimize compression costs, the hydrogen is produced from the
hydrocarbon feedstocks by reforming or partial oxidation at elevated pressures
of, e.g., from 100 to 1500 psig (pounds per square inch gage). Reforming or ;~
partial oxidatlon produces a mixture of hydrogen and carbon monoxide which is ~`

then treated to convert the carbon mono~ide to hydrogen and C02 by the so-
called shift reaction after which the C02 is removed prior to ammonia `
synthesis. Depending upon the process route chosen (i.e., reforming or partial ;~
oxidation), the re~uired nitrogen is added to the hydrogen either prior to ~ ~;
carbon monoxlde shift and C02 removal or following these steps. ~sing reform~
lng, the nitrogen is introduced by adding air during the reformlng operatlon, ii~ ~
- ~ ; ~''.`'
whereas in partial oxidation tralns, the nitrogen is added to the hydrogen

following C02 removal and prior to compression to ammonia synthesis pressures. ;~

~ ~ In sach ammonia synthesis trains, the final production cost of the

9~ synthetic ammonia is almost entirely a function of the capital costs of the

;~ 20 plant and the cost of the hydrocarbon feed. Both of these factors are ~
- ,
importantly affected by the capital cost and the efficlency of the systems used
for shlt conversion of carbon monoxide to hydrogen and the removal of carbon
dioxide. This invention is concerned with improvements in the a~rangement of ~
th~~carbon monoxide shift system and the C02 removal system and the manner in ~ -
which thesa systems are integrated so as to provide an integrated system which
is low in capital cost and high in thermal efficiency and which minimizes
process gas losses in the ammonia synthesis train.
, . ,:
Currently, the most commonly used ammonia synthesls process employs ~ ~
: ~ :
gaseous or light liquid hydrocarbons as feedstocks and steam reforming process-

ing. The hydrocarbon feed together with steam is treated over a reforming ;~
... : : i ~ :


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catalyst at temperatures generally ranging from 1000F to lSOQ F and
pressures generally ranging from lO0 to 600 psig. The nitrogen in the
proper ratio required ln the ammonia synthesis reaction is added in the form
of air in a secondary reforming step. The carbon monoxide-hydrogen mixture
produced by reforming is then treated in a series of so-called shift conversion
steps to convert the carbon monoxide to hydrogen and C02. The process stream,
now consisting essentially of hydrogen, nitrogen and C02, is treated for ~he
removal of C02 following which small res1dual amounts of C0 and C02 are ;~
removed usually by a metha~ation step which converts the residual C0 and C02
into methane by reaction with hydrogen contained in the gas over a methanation ; -
catalyst. The process gas is then cooled and compressed to ammonia synthesis
s prelssures of, e,g., 2000 to 8000 psig.
After contact with the ammonia synthesis catalyst, ammonia is recovered
from the process gas and unconverted synthesis gas is recycled in the so-called
ammonia recycle loop to be retreated over the ammonia synthesis catalyst.
Because o this recycle operation, lnert materials (mainly methane and argon)
tend to build up in the recycle loop. In order to maintain the concentration
of these inerts at a reasonable level in the recycle loop, it is necessary ~ ;~
¦ to continuously purge a portion of the recycled gas from the loop at a ratel 20 which will keep the inerts at a constant tolerable level. The purge gas consists
,
l, mainly of hydrogen and nitrogen together with some unrecovered ammonia and;~
these inerts, and has little value except as a waste fuel gas. The loss `
, of hydrogen and ammonia in the recycle loop purge gas can represent a sub-
3 stantial loss of the total potential yield of ammonia.
;~ In ammonia trains based on the steam reforming of gaseous
; or light liquid hydrocarbon feedstocks, the most common procedure is to usetwo stages of carbon monoxide shift conversion in order to reduce the
'~ residual carbon monoxide in the synthesis gas to a relatively low level. In
j the first stage, the bulk of the carbon monoxide is converted in a so- ` -
called bigh temperature shift converter normally employing an iron oxide
¦ catalyst promoted with small amounts of another metal oxide such as chromium
~
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oxide at tempera~ures generally ln the range of from 600 F to 1000 F. The
high te~perature employed in the first shift onversion stage favors rapid
reaction and thus minimizes the amount of catalyst required. On the other
hand, the shift equilibrium constant, K for the reversible shift reaction~

, , ' ~ ~`;,.
CO ~ H O ~CO + ~
2 ~ 2 2 ;
is less favorable at higher temperatures. The equili~rium constant K for

the above shift reaction may be expressed as:
Kp = ~H2~ (C02)
tH20~ ~CO)
; where the parentheses indicate the partial pressure of the component designated
10 within the parentheses at equilibrium. At high temperatures, the values for
K are relatively low, reflecting relatively lower degrees of conversion of
carbon monoxide. Typically residual carbon monoxide content in high temperature
5hift effluent may be in the range of from 2.5% to 3.5% by volume (anhydrous `~
basis).
Since it is uneconomic to leave relatlvely high concentrations of
unconverted carbon monoxide in the synthesis gas (e.g., 2.5% - 305%), a second
l ~ shift conversion stage is employed using a so-called low temperature shift -~
-' catalyst, typically a reduced copper catalyst promoted with zinc oxide and
generally operating at temperatures in the range of from 350 to about 550F.
In the second stage of conversion, the shift equilibrium Kp i9 much more
favorable and typically, the carbon monoxide can be reduced to residual levels
of the order of 0.2% - 0.5% by volume (anhydrous basis) leaving the low
temperature shift cQnversion stage.
The process gas, after having thils been treated successively in high
temperature and low temperature shift conversion stages, now contains
typloally from 16% to about 25% C02 by volume (an~hydrous basis). C02 removal
is then carried out, usually by scrubbing with an alkaline liquid scrubbing -~
: :~ ~ . -

,~ a8ent ~uc~ 8s an aqueous solution of potassium carbonate or of an

ethanolamine to produce a gas containing generally less than 0.2% C02. The
;,
~ 3

f
~;

process gas now containing typically .3% carbon monoxide and .~ or less
of carbon diox~de is then further treated over a methanation catalyst to convert
the residual C0 and C02 eo methane by ~he following series of reactions~


C0 + 3H2 > CH4 ~ H2
.
C2 ~ 4~2 --~ CH4 ~ 2 2 ;~

Note that for each mole of C0 converted to methane, three moles of hydrogen are
consumed and for every mole of C02 methanated, four moles of hydrogen are
consumed. The hydrogen for the methanation reaction Ls of course supplied by
the hydrogen in the process stream. For example, in a typical case where
0.3% C0 and 0.2% C02 is methanated, there is a loss of approximately 2.3% of
the total hydrogen in the synthesis gas. Even more significant than this ~`
substantial hydrogen 1099 in the methanation step, the conversion of the
re~idualmounts of C0 and C02 to methane substantially increases the inerts
contents o the synthesis gas and very subs,antially raises the purge losses
of hydrogen and ammonia in the recycle loop of the ammonia synthesis reactor. ~ ;
In the typical case discussed above, the methanation of the residual C0 and C02
would introduce approximately 0.5% methane into the ammonia synthesis gas. ;
`~ This increased inerts content in a typical ammonia synthesis loop would result
1 in hydro8en 1099 in the purge gas equal to approximately 5% of the total
I 20 hydrogen: In most typical ammonia tra:Lns in current operation, losses of
¦ ammonia production due to these losses, i.e. loss of hydrogen in the
methanation step, and loss of hydrogen and a~nmonia in the purge step, may
~ run up to 10% or more of total ammonia production, equivalent in a large plant
I to losses of several million dollars a year in ammonia product.
There have been numerous suggestions for improving the efficiency of
ammonia synthesis trains ~y introducing a C02 removal step between shift
conversion stages in order to reduce the concentration of C02 entering the
fi~al shift conversion stage to~a low level in order to achieve a high degree ~ ~ `
of conversion of the C0 in the final shift. By reducing the concentration of

- : .
~ 30 C02 in the final shift conversion stage, more complete conversion of the C0 is

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~13859B ;
of course favore~ because of the reversible nature of the shift reaction.
See, for example, U.S. Patent No. 2,487,981 to Reed, No. 3,382,045 to
Habermehl et al, and No. 3,57~,221 to Smith et al. While a small number
of commercial plants have employed a C02 removal step between stages of shift
conversion, only minor advantages have been obtained in contrast to the
~` increased cost and complexity of the C02 removal system, the necessary -
addi~ional heat exchange equipment required, and the additional heat losses
. , .
incurred as the process stream is alternately cooled and heated between

successive shift conversion and C02 removal stages. As a cons~quence, the
., .~ .
great màjority of ammonia trains do not use between-shift C02 removal systems.

7 GENERAL DESCRIPTION OF INVENTION ~i
i AND PREFERRED EMBODIMENTS ;~

In accordance with the present invention a markedly improved arrange-

ment and integration of the carbon monoxide shift system and the C02 removal ~
.;j ~ .
system has been found which provides a sharp reduction in process gas losses,
and which at the sa~e tlme ls low in capital cost and hlgh in thermal
3 efficlency. Generally stated, the lnvention involves the following sequence
-~ of process steps. The s~lperatmospheric pressure mixture of hydrogen and carbon
monoxide generated 4y reforming or partial oxidation of hydrocarbons is
.' . . - '~
sub~ected to shift conversion in a series of at least two shift conversion
stages, the first of which ls a so-called high temperature shift conversion, ~ ;
., and the last of which is a so-called low temperature shift conversion. In
this series of shift conversion stages all but a small amount, viz. from 90
to 99% and preferably from 95~ to 98% of the total carbon monoxide in the gas
is converted to hydrogen and C02. There is no C02 removal between these shift
conversion stages.
~: ,
s The process gas leaving the low temperature shift conversion stage,
typlcally at a temperature of, e.g. 440F, is then cooled by heat exchange and

the thus cooled process gas is then sub~ected to a first stage of C02 removal
uBlng an aqueous solution of potassium carbonate wherein at least the ma~or
portion of the absorption is carried out at an absorption temperature in

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: , 1 . : : . . - : . : . ,: : ,

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the vicinity of the atmospheric boiling temperature of the solutlon
(typically 230 F), and wherein the regeneration of the potassium carbonate
scrubbing solutlon is carrled out at approxlmately atmospheric pressure by ~ -
steam stripping. In this first essentially isothermal and relatively high
temperature CO2 removal step, the bulk of the total C02 generated in the
a~monia synthesis train is removed with high thermal efficiency. The C02
content of the process gas is reduced to a residual level of not less than
0.1% by volume (anhydrous basis~ and not more than 2% by volume (anhydrous
basis) and preferably to a residual level Of not less than 0.2% and not more
than 1%.
The process gas stream, now containing only a small residual amount of -~
carbon monoxide and a small residual amount of C02, and at a temperature of
typically `210F, is then heat exchanged to raise its temperature and sub~ected `
to a final stage of shift conversion using a low temperature shift converslon
catalyst to convert all but a smal~l amount oP the residual carbon monoxide
to hydrogen-and C02. This inal low temperature shift conversion should be
carried out to produce a final treated gas containing not more than 0.1% and
preferably not more than 0.05j% by volume (anhydrous basis) of carbon monoxide. ;
The gas stream leaving the final low temperature shift, typically at a
1 20 temperature of 410F, is then cooled by heat exchange and the thus cooled
~ process gas, now containing only a small amount of C02, is sub~ected to a
! second stage of C02 removal using an aqueous alkaline scrubbing solution,
preferably an ethanolamine, in a nonisothermal cycle wherein the scrubbing
is carried out at a relatively low temperature, typically 120 F, with the
, regeneration occurring by steam stripping at about the solution boiling
temperature. While the thermal efficiency of the second C02 removal stage is
, much lower than the first stage, in the second stage the C02 is effectively
'~ reduced to very low values of not more than 200 ppm (parts per ~illion) by `~
volume and preferably not more than 100 ppm (anhydrous basis). Since only a "~
30 ml~ p~ffpffrtion of the C~2 is removed in the second stage and because the
7, permissible exit CO2 fr~m the hot potassium carbonate scrubbing system is



-6-


7 ,

10385Y~
greater than allowable if it were the only C02 removal system in the proceas
~:, :,.. ..
train, the average thermal efficiency of the two C02 removal stage~ i9 higher
than would be possible using either sys~em alone. The gas stream leavlng
the second C02 removal stage, now containing very small residual amounts of
CO and C02, is then subJected to methanation in conv~ntional fashion to
convert such small residual amounts to methane, after which the gas is cooled,
~ compressed and converted to ammonia over an ammonia synthesis catalyst in
l normal fashion, including a recycle loop with provision for the purging of ;
inerts. As will be illustrated by the example which follows, the reduction
~, . ~ . , . ~ . .
of hydrogen losses in the methanation step~ and in the purge from the~recycle
~ i . .
loop ac~ieved by this sequence of steps permits increases in ammonia
7' ~ ' production of the order of lOX from a given quantity of hydrocarbon feed
;~ compared to current practice.
In order to achieve these highly important savings without sufering ; ~;
offsetting disadvantages of lowered thermal efficiency and substantially
increased capital costs for the CO shift and C02 removal systems, it is
essential to adhere to~the sequence-and choice of the shift converslon and
C2 removal systems as~described herein.
~ It is.essential in the practice of the invention that the first series
;~ ~ 20 of shift conversion steps are carried out such that from 90~ to 99%, and
preferably from 95% to 98% of the carbon monoxide content of the process gas ls ~
converted to hydrogen and C02. This is critical for the following reasons. ~ ~ -
Firstly, the CO left~ in the gas after the first series of shift conversions
must be shited in the final low temperature shift converter. The higher
~ the~concentration of CO entering the final shlft converter, the higher the
,`t;,~ , concentratloD of C02 gsnsra~ed during ths flnal shlft, snd the pressncs of
such G02 has an increasingly advsrae sffect on the complsteness o~ ehe
reversible CO conversion reaction. Secondly, the C02 generated during the
finsl shift conversion~must be removed in ths second C02 removal stage, which
30 ~ i~ th~r~ally~inefficient compared to the much higher thermal efficiency of
the first isothermal C02 removal s~age, thus adversely affecting the overall



7-
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.: :, , , . , , , . . , :

thermal efficiency of CO2 removal. Thirdly, if the C0 concentration in the
gas entering the final low temperature shift conversion stage i8 high, then
additional st~am, not otherwise rèquired, must be added to the process gas
stream to supply that necessary to react with and comvert such C0 to H2 and

C2
The use of the essentially isothermal, relatively high temperature
C2 removal system employing aqueous hot potassium carbonate as the
absorbent to absorb all but a minor portion of the total C02 from the process
, gas between two stages of low temperature shift conversion, cGupledwith the
^` lO use of a non-isothermal scrubbing system for final C02 cleanup is similaFly
- important. Because the absorption temperature of the hot potassium
carbonate system approaches more closely the temperature of the low temperature
~i shift reaction, smaller heat losses are encountered as the gas stream is
alternately cooled and heated between the two low temperature shift stages
than if lower temperature absorbent systems are employed which generally
operate about at least 100F lower in temperature. Secondly, it has been
~, found that the capital cost and !thermal efficiency of the hot potassium carbon-
i~ ate absorption system reaches close to its optimum when operated to remove
C2 down to the levels required for close to optimum operation of the second
20 low temperature shift converter. At the same time, wlth the bulk of the ;
C2 removal duty being performed by the thermally efficient hot potassium
carbonate system, only small amounts of C02 remain to be absorbed in the final
. . ~ .
: non-isothermal scrubbing stage. The ability of the non-isothermal scrubbing
system to reduce the C02 to a final very low level more than compensates for
its lower thermal efficiency because of the sùbstantial reduction in hydrogen
losses in the methanator and recycle purge resuiting from the very low lPvel
of C2 in the gas entering the methanator.
~nother highly important aspect of the invention is the manner in which ~ ~-
~j the heat content in the process gasl both sensible and latent, is recovered ;~
and r~transferred to the process gas stream as the process gas leaving the
low temperàture shift converter is first cooled prior to entering the first
, ~




.. .

bulk C02 removal stage and then reheated following bulk C02 removal prior to -~
,
entering the final low temperature shift converter. In accordance with
the preferred embodiment of the invention, such heat transfer operation is
carrled out with a system that includes a series of direct contactors where
both the latent and sensible heat of the shifted gas i9 transferred to a
circulating stream of water which 8 then employed to reheat and resaturate
the process gas leaving the first stage of C02 removal before it enters the
final low temperature shift converter.

- ~:
In accordance with a particularly preferred embodiment of the invention, ~ ~
, ;. ~
10 the hot process stream leaving the first low temperature shift converter, is
contacted, prior to C02 removal, with a circulating stream of water in a
direct contactor, cooling the process gas and heating the water. The thus
cooled, C02 containing gas is then scrubbed in the first C02 scrubbing stage
to remove the bulk of C02 down to a resldual level of, e.g. 0.5 to 1%. The
thus treated process gas from the flrst C02 removal stage, now typically
reduced in volume by, e.g. 15% to 30% compared to the volume of the gas
prior to C02 removal, is then contacted in a direct contactor wlth the hot
~; water heated by the higher volume process gas stream prior to C02 removal.
j~ Both the se~sible heat of the CO2 content of the effluent from the first low
:i , .
20 temperature shift convertër, and the latent heat of the steam associated
with such C02 content, is transferred with a high degree of efficiency to ~
the lower volume process gas stream passing from the bulk C02 removal stage ~ ;
to the final low temperature shit conversion stage.
-l In accordance with preferred embodiments of the invention, the heat -~
J- required to generate the stripping steam for the first and second C02
removal stages is supplied by the heat content of the hot process gas. It
J is possible to supply the full steam stripping requirements of both CO2 .
~,~ removal systems entirely from the heat contained in the process gas by

virtue of the manner in which the C02 removal systems are integrated with the

shift conv~slon systems, enabling both systems to operate at a maximum
thermal efficiency.


. -i:



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"' '~ ' ' ' ' ', '

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Reference is now made to the accompanying drawing wherein the Flgure is -~
a schematic flowsheet showing an ammonia syn~hesis train wherein the CO shift
conversion systems and the CO2 remova~ systems have been integrated in
accordance with preferred embodiments of the invention.
In the embodiment shown ln the Figure, the hydrocarbon feedstock is a
gaseous or light liquid hydrocarbon which is introduced into the primary
reformer l through line la. Steam is introduced through line 2 and the
mixture flows into the reformer through line 3. In accordance with the usual

,
practice for ~he steam reforming of gaseous or light liquid hydrocarb ns
up to about the naphtha range, the mixture of steam and hydroca~bon is
passed through tubes b~ packed with a catalyst 5 which is normally a nickel
catalyst supported on an alumina base. The tubes 4 are externally heated by
burning a fuel-air mixture introduced by line 6 around the exterior of the
tubes such that the steam-hydrocarbon mixture is sub~ected to temperatures of
1000F to about 1500F as it passes over the nickel catalyst. In the reform- ;
ing reaction most oE the hydrocarbons are converted by reaction with steam
to a mixture of hydrogen and carb n monoxide and minor amounts of carbon
dioxide and unreacted hydrocarbons. Using methane as the hydrocarbon feed~
stock, for example, in normal practice about 70% of the methane will be
~ ~ '
reformed to a mixture of hydrogen and carbon monoxide and small amounts of
carbon dioxide such that the gas leaving ~he reformer by line 7 will contain
about 10% by volume of residual methane. i
~ The hot mixture leaving the primary reformer is conducted by line 7
i~ to ~econdary reformer 8, provided with a bed of a nickel reforming catalyst 8a.
Upstream from the bed of catalyst 8a, air is introduced through line 8b which ;
, burns a portion of the reformed gases raising the temperature of the
3~ process gas stream typically to a temperature of 2000 F. As the superheated
mixture passes through~the catalyst bed 8a, unreformed hydrocarbons are largely -
~! converted into carbon monoxide and hydrogen and the hot process gas stream, now
,
conta~nin~ nor~ally only a few tenths percent of unconverted hydrocarbons, is


withdrawn from the secondary reformer by line 9~ The amount of air introduced
:i
: ~10-- -
. .,

,. . . .
. .


. . .

i ~ 5 ~
through line 9 i9 chosen to provide the correct proportion of nitrogen for
the subsequent ammonla synthesis reactor.
The primary and secondary reformers are operated under ssubstantial
superatmospheric pressures generally ranging from lQ0 to 600 psig, such
superatmospheric pressures being required not only for the economical
operation of the reformer, but also for the economic operation of the shift
conversion and CO2 removal systems and to minimize the cost of compressing the
synthesis gas to ammonia synthesis pre~isures.
The superatmospheric pressure gas mixture lsaving the secondary reformer
lC5 by line 9 consists primarily of a mixture of steam, hydrogen, carbon monoxide,
nitrogen, CO2, and very small amounts of unconverted hydrocarbons. Using
'~ methane as a hydrocarbon feedstock, a typical gas leaving the secondary
reformer would conta:Ln 22% nitrogen, 55% hydrogen, 13% carbon monoxide, 8%
carbon dioxide, 0.3% methane, and 0.3~ argon (all on an anhydrous basis) together
with about 60~ H20 as steam based on the total dry gas components.
In order ~o effect the conversion of the carbon monoxide content of
.
the gas leavlng the secondary reformer to additional hydrogen, the gases pass
~i through a series of shift converters operated and integrated with the C02 '
removal system in accordance with the invention. The shift conversion system
~'~ 20 in the embodiment shown in the Figure consists of high temperature shift
conversion reactor 14, followed by primary low temperature shift con~7ersion
reactor 23, without intermediate CO2 removal, to effect the conversion of all
but a small residual amount of ~he carbon monoxide content of the gas.
~, Followingheatexchange and a first stage of CO2 removal as will be described
'1 ' , ,
, below, the gas is then subiected to a second and final stage of low
temperature shift conversion in secondary low temperature shift conversion
reactor 26 to reduce the C02 content to a very low level.
Attention is now directed to the first two stages of shift conversion
~ carried out in the high temperature shift conversion reactor 14 and in the; 30 primary lbw temperatureshift conversion reactor 23. The superatmospheric
j : ~
pressureprocess gas mixture leaving secondary rsformer 8 by line 9 typically
,



.. ,~: . ~ : -. , . .. . ,. ~ . .

1~3~59~
at a temperature of 1850F is first passed thro~lgh heat exchanger 10 where - `
; it passes over a tube bundle 11 supplied with boile~ feed water through line
lla and delivering steam through line llb. In normal pract~ce high pressure
steam a~, e.g. 1500 psig is raised in heat exchanger lO which is used to
operate steam turbines driving the synthesis gas compressors or other
` mechanical equipment required intheprocess.
; The thus cooled process gas stream now typically at a temperature of
, e.g. 700F is then introduced by line 13 into the high temperature shift `~
, . ~
conversion reactor 14 and passed over high temperature shift conversion
catalYst 15 at a temperature in the range of from 600-1000F and usually
in the range of from 650-900F. The high temperature shift conversion
catallyst is normally an iron oxide catalyst usually promoted with small
amounts of another metal oxide, such as chromium oxide. Suitable high
temFerature shift conversion catalysts are described, for example, in United
Stat-es Patents 2,364,562; 2,602,020; 2,631,086; and 2,815,331.
In the high temperature shift conversion reactor 14, because of the ;
relatively high temperature employed, the conversion of the carbon monoxide
- is incomplete for the reason explained, namely that the shift equilibrium `~
constant, Kp, for the,reversible shift reaction is less favorable at higher ~-
temperatures and the presence of the relatively large amount of shifted C0
urther reduces the amount that can be shifted at a given temperature. ;-
Generally, it is possible to convert only about 80~ of the carbon monoxide
in the~high temperature shift converter and typically the residual carbon ~ ;
monoxide content leaving the high temperature shift converter by line 16
~ is of the order of 2.5~ to 3.5~.
-1 The effluent from high temperature shi~t converter 14, leaving at a
~ temperature typically of e.g. 850F is conducted by line 16 to a heat exchanger
i-~ 17, where it-passes over a tube bundle 18, supplied with boiler feedwater
through line l9a and delivering steam through l9b. The cooled effluent,
typlcally at a temperatùre of e.g. 630F, is conducted by line 20 to a
gas-to-gas heat exchanger 21, where it passes over a tube bundle 77 carrying
,'i ::
-12-

,
::

:~ ~

:~03859~ ;
process gas which has been treated in the second C02 removal unit and which
requires preheating prior to methanation. l~e process gas from llne 20
flowing over the tube bundle 77 in heat exchanger 21 preheats the process
gas prior ~o methanation and is itself cooled, leaving heat exchanger 21 by
" , ,
line 22 typically at a temperature of 420F.
The thus cooled process gas stream ls then imtroduced into the primary
low temperature shift converter 23, and passed over low temperature shift
conversion catalyst 24. The low temperature shift conversion catalyst is ~ ;
normally a copper-zinc catalyst made by preparing a mixture of zinc and copper
oxide and activated by a controlled reduction of the copper oxide to metallic
copper. Suitable low temperature shift conversion catalysts are described,
s for example, in United States Patents 3,303,001; 3,390,102; 3,546,140; and
3,615,217.
These catalysts operate in the temperature range of Ero~ 350F to 550F '
` and usually in the range of from 400F to 500F.
In the low temperature shift conversion reactor 23, because of the
~ lower temperatures employed, the conversion of the carbon monoxide is
`, carried to a higher degree of compietion since the shift equilibrium constant
R for the reversible shift reaction is more fa~orable at the lower temperatures..
, 20 However, because of the presence of substantial quantities of C02 (usually
in the range of from 15% to 25%, anhydrous basis) it is generally not possible ;;
, ' ~
to reduce the concentration of C0 below about 0.2% in the presence of such
~j substantial amounts of C02 and typically the residual carbon monoxide content
leaving the primary low temperature shift converter 23 will range from 0.25%
. ~ .
to 0 35%
~ As previously pointed out, it is essential in the practice of thel invention that the con~erson of CO in the first series of shift conversion
`; f
reactors be carried to a 90-99%, and preferably 95-98% level of completeness.
~ Higher amounts of unconverted C0 leaving the primary low temperature shift
:! 30 conv~rsion not only adversely affect the performance of the secondary low
`1 temperature ~hift converter 26 in its ability to reduce the final C0 to low
values but also adversely affect the thermal efficiency of the overall C02
.
-13

'J '
. '
'.. : ' , :, , . , . . . .. ` : .

removal system. During the operation of the low temperature shift co~version
catalyst 24 the catalyst gradually declines in activity. This decline,
whlch is usually due to the loss of catalyst surface ~rea with time and/or
the poisoning effect of small traces of sulphur compounds in the process gas,
is evidenced by a rise in the operating te~perature required to maintain
catalyst activity with an attendant rise in the C0 concentration in the ;shifted gas. When the C0 concentration rises to an unacceptable level,the
catalyst is normally replaced with a new charge. The references herein to the
percent conversion of the C0 during the shift reaction or to the residual C0 ~;
in the shifted gas are intended to refer to the C0 values obtained during
the middle portion of the catalyst run rather than the values obtained
initially or at the end of the run when the catalyst must be replaced because
of loss oE activity.
The process gas stream leaves the primary low temperature shift converter ~`
23 by linë 25 typically at a temperature of 450F and containing approximately ;~
18% C02. It is then conducted to tube bundle 28a of gas-to-gas heat exchanger
28 where it preheats process gas enterlng heat exchanger 28 through line 57
which has been previously subjected to the first stage of C02 removal. The
process gas stream, typically at a temperature of 390F, is then conducted by
i 20 line 29 to direct contactor tower 30. In direct contactor tower 30, the ~ ~;
process gas entering at the bottom through line 29 is contacted with the
circulating water introduced into the top of the tower by line 32 and liquid
distributor 33. The tower 30 is equipped with tower packing or contact plates
indicàted by numeral 31 to insure intimate contact between the water flowing
.. . .
; down over the packing or plates 31 and the process gas stream rising upwardly -~
through the tower counter-currently to the descending liquid. - ?
ot water accumulating at the bottom of tower 30 in sump 34 is
.. ~ . .
conducted by line 35 t~ a second direct contactor tower 36 similarly equipped - ~ ~
:! ~
; with paeking or contact plates 37. The hot water èntering tower 36 by line 35
-~ 30 ~ dl~t~ibue~d by liquid distributor 38 over the packing or contact plates 37,
l and flows downwardly through the tower 36 counter-currently to process gas -
'. . ' : '' '
-14-
! :

,

:~ . .` . . . ` .

enterlng the bottom of the tower through line S6. ;~;
Water, cooled by contact with the process gas, collects at the bottom of
contact tower 36 in sump 39 and is returned to the top of contact tower 30 ~;
by lina 40, circulating pump 41, and line 32. As will be explained in more
detail hereafter, through ~he agency o~ the pair of dlrect contact towers
30 and 36 and the water that is circulated around these towers, heat contained
hn the relatively hot process gas stream entering tower 30 through line 29 ;
, i9 transferred to the relatively cool process gas stream entering tower 36
~ through line 56.
" .
The process gas stream after having given up a portion of lts heat
content in direct contact tower 30 leaves tower 30 by line 42 typically at a
temperature of 290 F and is conducted to reboiler 43. In reboiler 43j the
process gas passes through tube bundle 44 where it heats scrubbing solution in
~; the second, non-isothermal C02 removal unit as will be descrlbed subsequently.
j The process gas stream leaving tube bundle 44 typically at a temperature of
1 260F i8 then conducted by line 45 to knock-out pot 46, where steam condensate
::1 . , ,
collects in sump 47, a portion of the condensate being in~ected through line
48 into the process gas stream flowing in line 56 and another portion being
removed from the system by llne 49 for any desired use.
The process gas etream leaving the knock-out pot 46 by line 50 is
introduced into the bottom of absorber tower 51, wherein the bulk of the C02
content of the gas is removed by contact with a hot aqueous solution of
.1 . ...
¦ potassium carbonate as will be described below in more detail. The hot
potassium carbonate scrubbing system comprises absorber tower 51, provided
~ with tower packing or contact plates 53 to insure inti=ate contact between
i gas stream and the potassium carbonate scrubbing solution, and regeneratibn
..,
tower 52 supplied with tower packing or contact plates 54 to lnsure imtimate ~;
contact between stripping steam and the potassium carbonate solution. Other
¦ details of the hot potassium carbonate scrubbing solution will be described
.' i :
30 3ubse~uentl~.
After the removal of the bulk of the C02 from the gas stream in

~ -15-
. .
..~



:.: . . , . . : .~ :: : : : : .

absorber tower 51, the process gas stream leaves the top of the absorber
tower by line 55 typically at a temperature of 210 F and typically contalning
0.5% C02. The process gas stream then 1OWS by line 56 to direct contact
tower 36 where it is brought into direct contact with water that has been ;
heated to a temperature typically of 340F by the hot process gas stream
entering direct contact tower 30 by line 29. After passing through direct
contact tower 36, the process gas is removed at the top of the tower by line
57 typically àt a temperature of 330F and saturated with water. The process ~j
gas stream then passes through gas-to-gas hea~ exchanger 28 where it flows over
., . . :; :~,~
tube bundle 28a and is heated by the process gas coming from primary low
... ~ ,: ~ . .
temperature shift converter 23 by line 25.
The process gas stream, now typically at a temperature of 400F flows
by line 58 into secondary low temperature shift converter 26 provided with
low temperature shift conversion catalyst 27 where the small amount of C0

....
remaining in the process gas after prlmary low temperature shift conversion

is sub~ected to further shift conversion with the result that the treated gas
.. . . .
leaving the secondary low temperature shift by line 59 contains only a very ~ ,
small residual amount of carbon monoxide typically in the range of from 0.01%
' to 0.03% (anhydrous basis) by volume.
The low temperature shift conversion catalyst 27, employed in the
secondary low temperature shift converter 26, will normally be a copper-zinc
catalyst of the same general type employed in the primary low temperature
shift converter 23, and will operate in generally the same temperature range,
i.e. from 350F to 550F and usually in the range of from 400F to 500 F.
It will often be advantageous to operate the secondary low temperature
shift converter at a somewhat lower temperature than the primary low
' temperature shift converter, e.g. from 50 to 75 F lower in temperature mainly
3 because the reduced C0 load to be shifted by the secondary low temperature
shift converter leads to a lower temperature rise through the shift catalyst bed.

The ability of the secondary low temperature shift converter to reduce
the carbon monoxide to very low residual levels results from a combination of
~ . -16-
~``,. ~.
!


the following factors. First~ because of the low shift temperature employed,
conversion of the carbon monoxide is favored because of the favorable shift
equilibrium constant K for the reversible shift reaction. Secondly, because
P
all but a small percentage of the C02 has been removed in the first C0
removal stage, the secondary low temperatur~ shift conversion takes place in
3 the presence of only a small amount of C02, which favors the more complete
~. .
conversion of the carbon monoxide to C02 and hydrogen. Thirdly, because the
amount of C0 entering the secondary low temperature shift converter is low,
only small amounts of C02 are formed during the shift reaction, which again
favors the more complete conversion of the C0. Fourthly, because of the small
amount of C0 entering the secondary low temperature shift converter, it is `
- . ~ '
possible to achieve a high steam to C0 ratlo in the secondary low temperatuse `~
shift conversion without incorporating large amounts of steam in the process
gas which would lower the thermal efflciency of the system. A high steam to
C0 ratio~ of course additionally favors the more complete conversion of the C0.
i Thè process gas stream leaving the secondary low temperature shift
conversion reactor 26 by line 59, typically at a temperature of 410~, and `~
containing e.g. 0.02% C0 and 0.7% CO2, is then conducted to reboiler 60.
In reboiler 60 the process gas passes through tube bundle 61 where it heats
the potassium carbonate scrubbing solution in the first essentially isothermal
C2 removal unit as will be descrIbed subsequently. The process gas stream
leaving tube bundle 61, typically at a temperature of 250F, is then conducted
J by line 62 to heat exchanger 63 which, in the embodiment shown ls a boiler
feed water heater containing tube bundle 64a supplied with boilèr feed
water through line 64b, heated boiler feed water being delivered through line ;~
64c. The process gas leaves boiler feed water 63 by line 65 typically at a
~, temperature of 120F and is conducted to knock-out pot 66 from which condensate ;
is removed by line 67 for any desired use.
The process gas stream is then conducted by line 68 to absorber tower 69
. '~ 30 of the ~econd non-isothermal C02 removal stage comprising absorber tower 69
provided with tower packing or contact plates 71 to insure inti~ate contact
, .... .
17
:''', ` '`' '`

~3~i~ ~between the gas and the scrubbing solution. Regeneration of the scrubbing ~ ;
solution occurs in stripping tower 70 provided with tower packing or contact
plates 72. The details of the secondnon-isothermal C02 scrubbing system will ~ ~-
be described subsequently.
In the second non-isothermal C02 removal unit, the relatively small
amount of CO2 in the gas entering absorber 69 is reduced to a low level,
typically of the order of 10-50 ppm and the process gas stream thus raduced in
C2 content leaves the top of absorber S9 by line 73 and is conducted to gas-to-
gas heat exchanger 74 where it flows over tube bundle 75 where it is heated
by relatively hot process gas flowing through tube bundle 75 coming from
methanation reactor 79. The process gas now typically at a temperature of
230F flows by line 76 to a second gas-to-gas heat exchanger 21 where it flows
through tu~e bundle 77 and is heated by relatively hot process gas entering
the shell of heat exchanger 2L from 1ine 20. The process gas stream, now
typically heated to a temperature of 600F, flows by line 78 to methanation
reactor 79 provided with a methanation catalyst 80. The methanation catalyst
80 is normally a nickel catalyst and operates at a temperature of generally `;
in the range from 550 F to 800F. In the methanation reactor 79 small amounts
. A . .
of residual C0 and C02 are converted to methane by reaction of the CO and C02
:~ .
20respectively with hydrogen present in the process gas stream such that the
combined residual concentration of C0 and C02 in the gas stream leaving the
methànation reactor by line 81 i8 generally less than 10 ppm. The process gas
stream is then conducted by line 81 to heat exchanger 82 where it flows
over tùbe bundle 82a supplied with water through line 82b, hot water or steam
leavlng by line 82c.
The process gas stream now typically at a temperature of 300 F is -~
;, conducted by line 83 to gas-to-gas heat exchanger 74 where it flows through
tube bundle 75 and heats the process gas stream flowing through the heat
exchanger from line 73 and itself is cooled. The process gas stream now
30typlcall~ at a temperature of 190F leaves heat exchanger 74 by line 84
and is conducted to cooler 85 and then is conducted by line 86 to knock-out

.','`' . '
-18-

. ,~,
,` .- ' , ' 1 ' . . . ' ", ~ . I . , , ~ .' !

1~3859~ ; ~
pot 87 ~yplcally at a temperature of 100~. Condensate separated in knock-out
pot 87 is removed by line 88 for any desired use and the cooled gas st~eam
is then conducted by line 89 to synthesis gas compressor unit 90 where the ga~
is compressed to ammonla synthesls pressures typically of the order of from 2000to 8000 psig.
After compresslon and introductlon of the recycle gas from line
96 the combined process gas stream ls then conducted by llne 91 to ammonia
synthesis converter 92 provided with ammonia synthesis catalyst 93. The ammonia
synthesis catalyst is normally a promoted iron catalyst and operates usually
at temperatures in the range of from 650 to 900 F. In the ammonia synthesis
converter, partial conversion of the nitrogen and hydrogen in the gas stream to
ammonia occurs, the usual percentage conversion per pass over the ammonia
synthesis catalyst being of the order of 15% to 30%.
The process ga~ stream leavlng the ammonla synthesis reactor by llne
9~ typically contains 10% ammonia, 10% nltrogen, 54% hydrogen and the balance
1s inerts consisting almost entirely of methane and argon.
The process gas stream then enters ammonia' recovery u~it 95 where
most of the ammonia ls recovered by chilling the process gas, ammonia product
being withdrawn by line 97 and unconverted hydrogen and nitrogen being recycled
to the synthesis gas compression unit 90 by line 96 for compression and
retreatment over ammonla synthesis catalyst 93.
The system comprising ammonia synthesis compression unit 90,
ammonia synthesis reactor 92, ammonia recovery unit 95, and recycle llne 96
is generally referred to as the ammonia recycle loop. In order` to prevent
t the accumulation of inert materials consiseing of methane and argon in this ~
ammonia loop, it is necessary to continuously withdra~ a purge stream 98, ~ ;
from the ammonia loop in order to maintain the concentration of these inert
materlals at a tolerable level. ~sually the concentration of inerts in the
ammonia loop is maintained at a level of from 15% to 30~. The composition of
this purge gas is generally equivalent to the composition of the recycled
gas stream 96, and consists mostly of nitrogen, hydrogen, a small amount of

-19-

.

~:

~C!38~9`~
ammonla and small amounts of methane and argon. Typically it wi}l contain
60% hydrogen and 3æ ammonla. It has little value except as a waste fuel
and represen~s one of the ma~or losses of hydrogen and ammonia product in the
entire system. As will be illustrated in the example which follows, the higher
the concentration of methane entering the ammonia loop in the synthesis gas,
the higher must be the rate of purge from the ammonia synthesis loop in order
to maintain the concPntration of lnert materials in the loop at a tolerable
level and, consequently, the greater the loss of the valuable hydrogen and
ammonia product in the purge gas. By operating in accordance with the invention,
the amount of methane enterlng the ammonia loop is greatly reduced, thus greatly
reducing the purge losses of hydrogen and ammonia product.
Attention is now directed to the first substantially isothermal C02
i scrubbing system including ab~orber tower 51 and regeneration tower 52. As
previously described the process gas enters the bottom of tower 51 thro~gh line
50 and after contact with hot aqueous potassium carbonate scrubbing solution is
recovered from the top of scrubbing tower 51 by iine 55. As the process gas
travels upwardly through the tower through section 53 provided with tower
packing or contact plates to insure intimate contact between the gas and
scrubbing solu~cion, it is brought into contact with a descending stream of
aqueous potassium carbonate scrubblng solution introduced into the top of
tower 51 through line 99 and liquid distributor 100. The scrubbing solution `
containing absorbed C02 collects at the bottom of tower 51 in sump 101 and is ~ ;
conducted by line 102, pressure letdown valve 103 and line 104 to the top of the
stripping tower 52 where it is distributed over the tower packing or contact
~ plates 54 by liquid distributor 105. As the solution passes through pressure
-~ letdown valve 103, the pressure is reduced from the superatmospheric pressure
prevalling in column 51 to the pressure prevailing in strippingtower 52, viz.,
atmospheric pressure or a pressure slightly abQve atmosph~eric. As a result
of the pressure reduction, a portion of the C02 is reieasedfrom the solution
~t t~ top o~ the stripping tower 52 and the partially regeneratedsolution then
flows downwardly through section 54 counter-currently to upwardly rising



- . :

~L~3~
3tripping steam generated at the bottom of the tower.
Stripping steam i9 provided at the bottom oE tower 52 partly by
means of solution reboiler 60 which is heated by hot process gas flowing
through the tube bundle 61. Solution collecting at the bottom af stripping
tower 52 in sump 106 is conducted by line 107 to the shell of reboiler 60
and then i8 returned together with steam generated in reboiler 60 to the
bottom of tower 52 by line 108. Additional stripping steam is al90 provided
to the bottom of tower 52 by line 109. As will be described more in detail
subsequently, the stripping steam delivered by line 109 is derived from the
second non-isothermal C02 scrubbing system.
As the scrubbing solutlon flows downwardly through section 54 of
tower 52, it is contacted with upwardly rising stripping steam resul~ing in
the desorption of the C02. The regenerated scrubbing solution collecting
~ at the bottom of tower 52 in sump 106 is withdrawn by line 110 and is then
3 returned by solution circulation pump lll and line 99 to the top of the
absorption tower 51.
~, The mixture of steam and C02 collecting at the top of stripping
tower 52 is removed from the top of the tower by line 112, pas~ea through
, condenser 11-3 and then is conducted by line 114 to knock-out pot 115.
Condensate collecting in knock-out pot 115 is returned by line 116 to the top ,
of stripping tower 52 in order to maintain the water'balance of the system.
C2 is removed from knock-out pot 115 'by line 117 for any desired use.
i '
The aqueous potassium carbonate solution empl,oyed in the first

,~ substantially isothermal C02 removal system is a relatively concentrated ,~

aqueous potassium carbonate solut-lon having potassium carbonate~concentrations

, by weight of from 15% to 45% and preferably from aboue 22% to 35% (these
.... ~.- .
concentrations by weight being calculated on the assumption that all the

i potassium present is present as potassium carbonate). Such patassium'

carbonate solutioDs are preferably activated by the addition of addltives such

~ 30 aa ~thanol8minss; alkali metal borates, such as potassium or sodium borate;
amino ~cids such as glycine; As203, or other additives which tend to increase


-21-

:` :

~ 38S98 ~ `
the rates of absorption and desorption of C02 in ~he potassi~ carbonate
solution.
Particularly preferred among these activators are the alkanolamines
which are preferably added to the potassium carbonate solut~on in amounts ;~
ranging from about 1% to 10% by weight, and preferably from about 2æ to 6%
by weight. Viethanolamine is preferred from the standpoints o~ cost, ` ~`
relati~ely low volatility, and effectiveness. However, monoethanolamine, or
triethanolamine, may also be employed in place of diethanolamine or mixtures ~; ?~
of any 2 or 3 of these ethanolamines, may be employed. ~ `
As is well known, the absorption of C02 in the potassium carbonate
solution produces potassium blcarbonate while regeneration or desorption releases
C02, producing potassium carbonate. The reversible absorption and desorption
i reactions do not go to completion in either the absorption or regeneration stages,
i ~nd consequently, the scrubbing solution i9 as circulated actually a mixture.
~ The regenerated scrubbing solution Eed to the absorber is a carbonate-bi-
J. carbonate mixture rich in carbonate while the solution leaving the absorber is ~ ,
a mixture rich in bicarbonate. References herein to scrubbing solution of
potassium carbonate are o~f course intended to include mixtures of potassium ~ -
,, ( ,
carbonate with potassium bicarbonate formed during the absorption process.
As pointed out previously, the first C02 removal system is substantially
isothermal in that the absorption and regeneration stages are both carried out
substantially at temperatures in the vicinity of the atmospheric boiling temper~ature of the solution. As is appsrent from the embodiment shown in the Figure,
, there is no heating or cooling of the~solution as it is continuously circulated ~ ~
`~i between the absorption and regeneFation towers. In some cases, it may be ; ~ ~;
desirable to cool a minor portion of the regenerated scrubblng solution before ~
. ~ ,
recycle to the absorption tower and to deliver this minor portion to the top of
the absorption tower while the bulk of the uncooled solution is delivered to
~he absorption tower at an intermediate point. In some cases, such minor cool~
30 ing of the ~tream delivered to the top of the absorption tower may be employed
to achieve a further reduction in the resldual concentration of C02 in the
,',' '~
-22-
',.: , : ,

~38S9~ ~
scrubbed gas leaving th~ absorptlon tower without substantially affecting the ~;
thermal efficiency of the system. In any event, at least the maior portion
of the absorption in tower 51 should occur at a temperature in the viclnity ~ ~
of the aS~mospheric boiling temperature of the solutlon, and the mean absorption ~ ~ ;
temperature prevailing in the absorption tower will generally range from 210 F
to 270F and preferably from 220F to 250F.
While the absorption in tower 51 takes place at the superatmospheric
pressures prevailing in the gas ~ynthesis train, the regeneration of the
soluSion in stripping tower 52 i8 carried out at pressures at or close to
atmospheric pressure. The stripping tower 52 would generally be operated at
pressures ranging from atmospheric to about 35 pounds per square inch gage,
whereas the pressures prevailing in absorption tower 51 will generally be in
excess of 150 pounds per square inch gage and usually in the range of from 250
' to 1500 per square inch gage.
In accordance with the invention the C02 content of the process gas
is reduced to a residual level of not less than 0.1% by volume and not more than
2% by volume in the first substantially isothermal C02 removal unit employing
hot potassium carbonate as the absorbent. It has been found that if the C02
removal in the hot potasslum carbonate unit is carried below a resi~ual C02
level of less than 0.1%, the overall thermal efficiency of C02 removal in the
. ~
system as a whole (i.e., the system including both the hot potassium carbonate
unit and the subsequent non-isothermal C02 scrubbing unit) declines
substantially because of the increased requirements for re~eneration steam in
the hot potassium carbonate unit, and at the same time the cost of the hot
potassium carbonate scrubbing unit increases substantially. On the other hand,
it has been found that if the C02 content of the process gas is not reduced in ~ -
the hot potassium carbonatè unit\to at least a level of 2% by volume, the overall
thermal efficiency of the system as a whole declines substantially because of
the increased C02 removal duty imposed on the subsequent non-isothermal
~crubblng unit with little or no offsetting increase in the thermal efficiency
~3 of the hot potassium carbonate scrubbing unit. At the same time, if the C02

., : ,
~ -23- -
,."~, , ' '' ' .


1~3~9~
content of the process gas leaving the hot potassium carbonate unit becomes
more than 2%, the operatlon of the secondary low temperature shift converter i8
adversely affected in that it becomes difficult to achieve the desired
conversion of the C02 content of the process gas to low levels. Preferably,
the C02 content of the process gas is reduced in the hot potasslum carbonate
unIt to a residual level of not less than 0.2% and not more than 1% by volume.
Atten~ion is now directed to the second non-isothermal C02 scrubbing
system including absorber tower 69 and regeneration tower 70. As previously
described, the process gas enters the bottom of tower 69 through line 68,
and after contact with the relatively cool scrubbing solution, it is recovered
from the top of the absorption tower 69 by line 73. As the solution travels ;~
upwardly through the tower through section 71 provided with tower packing
or contact plates to insure intimate contact between the gas and the scrubbing
solution, it is brought into contac~ wlth a descending stream of a relatively
cool scrubbing solutlon introduced into the top of the tower 69 through line
117 and liquid distributor 118. The scrubbing solution containing absorbed C02
collects at the bottom of tower 69 in sump 119. It is then conducted by
line 120 to tube bundle 121 of solution-to-solution heat exchanger 122 where it ;~
is heated by indirect heat exchange with hot regenerated solution. The heated
solution is then conducted by line 123, pressure letdown valve 124 and line 125
to the top o stripping ~ower 70 where it is distributed over the tower
packing or contsct plates 72 by liquid distributor 126. As the solution
passes through pressure letdown valve 124, the pressure is reduced from the
¢ . ~ , " ~,, .
superatmospheric pressure prevailing in column 69 to the pressure prevailing
in stripping tower 70, viz., atmospheric pressure or pressure slightly above
atmospheric. The solution in the stripping tower flows downwardly through
section 72 counter-currently to upwardly rising stripping steam genera~ed in
., .
the bottom of the tower.
Stripping steam is provided at the bottom of tower 70 by means of
~olutlcn reboiler 43 which is heated by hot process gas flowing through tube ~;
bundle 44. Solution collecting at the bottom of stripping tower 70 in sump 127
. is conducted by line 128 to the shell of reboiler 43 and then is returned ;
-24-

', . .

i9~ :
together with steam generated in reboiler 43 to the bottom ef tower 70 by line
129,
As the scrubbing solution flows downwardly through section 72 of
tower 70, it is contacted with upwardly rising stripping steam resulting in the
desorption of C02. The regenerated scrubbing solution collecting at the bottom
of tower 70 in sump 127 i8 withdrawn by line 130 and then conducted by solution
recycle pump 131, line 132, solution-to-solutlon heat exchanger 122, line 132,
cooler 133 and line 117 to the top of absorption tower 69. As the solution
travals from the bottom of regeneration tower 70 to the top of absorber tower
69, it is first coo1ed in heat exchanger 122 by indirect heat exchange with
relatively cool scrubbing solution passing through tube bundle 121 and then is
further cooled typically to a temperature of 110F in cooler 133 before being
lntroduced into the top of tower 69.
In the preferred embodiment shown in Figure 1, the mixture of steam
and desorbed gas collecting at the top of stripping tower 70 is conducted by
line 109 and line 108 to the bottom of stripping tower 52. To facilitate the

.j
transfer of steam from tower 70 to tower 52, the pressure in tower 70 is main- ` -
tained slightly higher (e.g. 0.2 to 3.0 psig higher) than the pressure in tower
52. The str-ipping steam is then reused in tower 52 for stripping the aqueous
potassium carbonate absorbent in column 52. Although the stripping steam leav-
ing tower 70 conta-lns some C02, the concentration of C02 is relatlvely small
because of the relatively small amounts of C02 removed in the second non-
isothermal scrubbing system. T~pically the gas entering the scrubbing tower
69 by line 68 may contain approximately 1% C02. The concentration of C02

:. :
~ in the stripping steam leaving the top of tower 70 is also kept low by `
`;~ employing a high stripping steam rate in tower 70. While ordinarily it would
be thermally inefficient to employ a high stripping steam rate in tower 70,
.,.j . ...
~, the overall thermal efficiency of the system is not affected if the stripping ~ ~
" !
, steam from tower 70 is reused in stripping tower 52. The high rate of


stripping ~t~am employed in tower 70 furthermore makes possible the very

, thorough regeneration of the scrubbing solution so that the residual C02 can


:' -25-
,:..:~j , ,


, ~ -... . .. : ~ .



be reduced to minimal levels in absorber tower 69.
Ihe scrubbing solution employed in the second non iso~hermal C02
removal system i5 an aqueous solution of an alkaline chemical absorbent capable .
of reacting with C02 to form a C02 reaction product which is capable of
regeneration by steam stripping causing the C02 to be desorbed. Typical
scrubbing solutions suitable for the second non-isothermal C02 removal stage
are alkanolamines, particularly the ethanolamines, such as monoethanolamine,
diethanolamine and triethanolamine. Such ethanolamines are normally employed
in aqueous solutions containing from 15% to 35% by weight of the ethanolamine.
10 Other suitable amine-type chemical absorbents are aqueous solutions of glycolamines such as diglycol amine, alkyl and cycloalkyl amlnes, polyamines, bi-
cycloamines and aromatlc heterocyclic amines. Inorganic chem-lcal absorbents mayalso be e~ployed ln the second non-isother~,al scrubbing stage such as aqueous `;
solutions of potassium carbonate or aqueous solutions containing mixtures
. .
of potassium carbonate and ethanolamines.
i Mean absorption temperatures in thesecond non-isothermal scrubbing
stage will be substantially below the atmospheric boiling temperature of the ~ -~
solution and will normally range from 90F to 140F and preferably from 100~
to 130F, while regeneration of the solution is carried out at the solution ;
20 boiling temperatures, namely temperatures in the range of from 220 to 260F.In contrast to the first C02 scrubbing st~,ge which is essentially isothermal
in character, the absorption and regeneration in the second C02 scrubbing
stage occur at substantially different temperatures such that a solution-to-
3 solution hea~ exchanger is required between the absorption and regeneration
stages to heat the scrubbing solution as it travels from the absorption tower tothe stripping tower and conversely to cool the regenerated solution as it travels
, from the stripping tower to the absorption tower. Because of the non-
isothermal character of the second C02 scrubbing stage, its thermal efficiency
is much lower than that of the first isothermal C02 scrubbing stage. However,
30 ~c~u~e only a small fraction of the C02 will normally be removed in the second
non-iso~hermal stage~ the overall thermal efficiency of the C~2 scrubbing
. .,
-26-
., ' .
. ~ .

S9~
system as a whole is not greatly affected. In accordance with the lnventlon,
; in a great ma~orlty of cases, not more than 15% and preferably not more than 5%
of the total C02 scrubbed will be removed by the second non-isothermal scrubblngstage.
In the second non-isothermal scrubbing stage, the concentration of
C2 in the process gas is reduced to a level of not more than 200 ppm and
preferably not more than 100 ppm. Because of the chemical nature of the
absorbent and the relatively low absorption temperatures, the C02 may be
reduced to these low levels quite readily in the second non-isothermal
scrubbing stage. As a result of such low CO2 levels, the losses of hydrogen
in the subsequent methanation step and the losses of hydrogen and ammonia in
the purge gas and the subsequent ammonla loop are greatly reduced.
E X A M P L E
The followlng example illustrates a speci~ic application of the
i invention using a natural gas a8 ~he Eeedstock for the ammonia synthesis train
3 and employing the process sequence illustrated i~ the drawing. Natural gas, ~ ~-
consisting essentially of =ethane, and superheated steam àra mlxed and preheated ;-
to 950F and introduced into a primary reformer containing a nickel catalyst
supported on an alumina base. The steam to carbon molar ratio in the feedstream
to the primary reformer is 4Ø The primary reformer i9 operated at a
' temperature of 1500F and at a pressure of 450 psig. The reformed gas from
the primary reformer ls then introduced into the secondary reformer where it
is mixed with air in amount to provide the correct ultimate nitrogen to hydrogenratio in the ammonia synthesis gas. The gas leaves the seconda~y reformer
at a temperature of 1830F and under a pressure of 410 psig. The hot mixture -
! is then passed ~hrough a heat exchanger where the gas is cooled to a temperature ~ ;
of 700F and the heat recovered generates high pressure steam which is employed
1 to drive steam turbines operating the synthesis gas com*ressor.
1 The process gas stream now at a temperature of 700F and a pressure
-1 30 of 400 p9ig~ and having the composition shown in column 1 of Table I, is
conducted to the high temperature shift converter where it is passed into contact

-~7-

.''

;', 1~38Sg~ .,.,.-
~ with an iron oxide high temperature shift conversion catalyst entering the :
-~ catalyst at 700F and leaving the catalyst at 810F. The shifted gas,
. leaving the high temperature shift convèrter, has a composition,as shown in ; "~
.~ , ,:
,'~ Column 2 of Table T. ' ~ ,'
Table I -:

Column 1 Column 2 Column 3
:, Component Gas ~ompositionGas Composition Gas Composition
'' After SecondaryAfter High Tem- After Prlmary Low , ~ :
:' Reformer Vol % perature Shift Temperature Shift .~
.5 10 - - - Conversion VoI ~ Conversion Vol % ,;~-.;., .
;, N2(Dry basis)* 22.80 20.80 20.30 '~
H2 " "55.6059.50 60.50 '~
'' CH " "0.270.24 0.24
'~ A " "0.280.26 0.25
~ C2 ~' "8.0516.10 18.23 .
; C0 " ~i 13.00 3.10 0.48 ~; ';~, .,
~, Total~' "100.00100.00 lO0.00 ;: ;;
Water content 60.3 42.2 , 4,2.2 '~
' Vol % as a
:":~ 20 percent of
~, Tctal Dry Gas
.~ * Dry basis = based on total gas excludlng water
','~ The process gas stream is then cooled from 810F to 440F while the ~ ~ :
',~ heat removed from the gas stream is used for the generation of steam.
::l The process gas stream now at a temperature of 410 F and at a pressure :-~
,' of 398 p8ig iS introduced into the primary low shift converter and contacted
;'! with a zinc-copper low temperature shlft conversion catalyst whereupon most of
thg,remaining C0 in the process gas is converted to hydrogen and C02. The ~ "' :.
~', process gas leaving the primary low temperature shift converter has a
composition as shown in Column 3 of Table l and i9 at a temperature of 440F
., 1 . . .
'~.`,,~ and at a pressure of 395 psig. ,
,-~, It will be noted that the carbon:monoxide content of the gas from the , :
~ secondary reformer (viz., 13.0%) is converted successively in the high ,,
,.,,"î temperature shift converter to a residual content of 3.1% and in the primary low :
temperature shift converter to a concentration of 0.48%. Thus, in the first
-,,'~ two stages o shift con~ersion 96.3% of the C0 is converted to hydrogen and C02.
.,,~, `
";.~, ' ;~
28- . :
- :
... .

~ B59~ ~:
It is to be noted that the concentration of C02 in the exit from the primary
~ow temperature shift converter is 18.2%. At this relat~vely high
concentration of C02, the shlft conversion of the C0 to very low residual
levels in the primary low temperature shift converter is not feasible.
The process gas Prom the primary low temperature shift converter
is conducted to a gas-to-gas heat exchanger (heat exchanger 28 in the Figure)
where it is brought into indirect heat exchange with process gas entering heat
exchanger 28 by line 57 at a temperature of 330F. The process gas stream
leaves the heat exchanger by line 29 at a temperature of 388F and is
10 conducted to direct contactor 30 supplied with water by line 32 at a ;~
temperature of 274F. While passing through the contactor 30, the process
gas is cooled to 291F and the water is heated to 337F and is conducted by
j line 35 to direct contactor 36.
The process gas i5 next conducted to the reboiler ~3 of the non-
isothermal C02 scrubbing unit, entering reboiler 43 at temperature of 291F
and leaving at a temperature of 260 F. At a temperature of 260F and at a
1 pressure of 390 psig, the process gas enters scrubbing tower 51 oP the hot
i potassium c~rbonate scrubbing system where it is contacted with an aqueous
I potassium carbonate solution containing 30~ by weight of potassium carbonate
¦ 20 and 3~ by weight of diethanolamine. In the hot potassium carbonate scrubbing
tower 51 the bulk of the C02 content oE the gas is removed and the gas leaves
the top of absorption tower 51 at a temperature of 210~F and has the
composition shown in Column 1 of Table II. As shown in Table II, the CO2
content of the gas leaving the potassium carbonate scrubber is 0.5% by volume
(dry ba~is).
.~ .




~29-

~o~
TAsLE I~
Column 1 Column 2 Column 3
Component Gas Composition Gas Composition Gas Composition
After Hot Potassium After Secondary After Non-Iso~
Carbonate C0 Removal Low Temperature thermal C0
i System Vol %2 Shift Removal System ;
Conversion Vol % Vol %
N2(Dry Basis) * 24.70 24.50 24.82
2 73.60 73.60 74.54
10C~4 " " 0.29 0.29 0.30
~ A " " 0.31 0.31 0.31
-1 2 0.50 1.28 30 ppm ;
;~ C0 " " 0.60 0.02 0.02 ; ~
Total" " 100.00 100.00 100.00 ~ ,
Water Content 2.8 33.4 .4
Vol % as a
~ percent of Total
} Dry Gas
¦ * Dry basis ~ based on total gas excluding water
'I 20 Following the hot potass1um carbonate scrubbing unit, the gas
,!, iS next condùcted at a temperature of 210F and at a pressure of 388 psig to
- direct contactor 36. In order to supply the water required for the secondary
low temperature shift conversion, a portion of the condensate from knock-out
pot 46 may be in~ected by line 48 into the process gas prior to introduction
into dlrect contactor 36. The gas entering direct contactor 36 by line 56 at -~
, a temperature of 210F contacts relatively hot water introduced into the top
of contact tower 36 through line 35 at a temperature of 337F.
, In the contact tower 36, the process gas stream is heated to a
temperature of 330F and is saturated with water as a result of the counter
-~i 30current contact between the gas stream and the hot water descending ~hrough
the tower. The heating of th& gas stream and its saturation wl~h water in
contact tower 36 is substantially facilitated by the fact that the volume of
process gas entering tower 36 by line 56 is substantially reduced in volume
with respect to the process gas entering contact tower 30 by line 29 as a
~ result ~ the fact that all but a small a~ount of the C02 (0.5%) has been
,

' ~ '
: ~:

~385~318
removed from the process ga~ in the hot potassium carbonate scrubbing system.
The process gas entering tower 36 by line 56 is only 82% of the volume of the
process gas entering contact tower 30 by line 29. This permits the higher
heat content of the higher volume gas stream entering contact tower 30 to be
- transferred to the stream of water that is circulated between towers 30 and 36
which in turn is transferred to the lower volume gas stream entering tower
36 by line 56. This in turn facilitates the reheating and resaturation of
the process gas following the first CO2 removal step to the temperature
required in the secondary low temperature shift conversion.
The gas stream now at a temperature of 330F and saturated with
water leaves the top of contactor 36 by line 57 and then is brought into
indirect heat exchange in gas-to-gas heat exchanger 28 with hot process gas
from primary low temperature shlft converter 23 and is heated to a
temperature of 400F an~ then is introduced into secondary low temperature
~hift converter 26 where it ls brought into contact with a zinc-copper low ~
temperature shift conversion catalyst where all but a very small residual port- `
. ion of the C0 is converted to hydrogen and C02. The process gas stream leaves
~, :
secondary low temperature shift converter 26 by line 59 at a temperature of ~;`
1 409F and has the composition shown in Column 2 of Table II. It is to be
-3 20 noted that the C0 is now reduced to a residual concentration of 0.02% (200 ppm)
while the C02 content is 1.28%. A portion of this CO2 was brought into the
gas by stripping of C02 from the water used in tower which i9 saturated with
C2 at a higher partial pressure in tower 30 than in tower 36.
~ The process gas stream leaving the low temperature shift converter
.~ 26 is conducted by line 59 to the reboiler 60 serv~ng the hot potassium
', carbonate scrubbing unit at a temperature of 409 F, leaving the reboiler 60
by line 62 at a temperature of 245F. It is then further cooled in boiler feed~
~, water heater 64 to a temperature of 120F and after passing through knock-out
i pot 66 is introduced into the bottom of absorber 69 of the second non-isothermal
.,
co~ ~2mo~1 gy8tem by line 6O where it is brought into contact with an aqueous
' solution of diethanolamine containing 22% by weight of diethanolamine.

~; -31-
. , .
.', ' - ;~
i'~;

1~38S9~
Regenerated diethanolamine enters the tap of absorber tower 69 by line 117 at
a temperature of 110F and the C02 contalning solution is withdra~m from the
bottom of tower 69 by line 120 at a temperature of 125F, passes through
solution-to-solution heat exchanger 122 leaving at a temperature of 210F, :~
and is then conducted by line 123, pressure letdown valve 124 and line 125
to the top of steam stripping tower 70 operated at a pressure of 10 psig where
,. .. ..
it is stripped with steam generated at the bottom of the tower. Regenerated ;~
solution leaves at the bottom of s~ripping tower 70 by line 130 at a ~ :
temperature of 245F and is then recirculated by pump 131, line 132, and
solution-to-solution heat exchanger 122 where it is cooled to temperature of .;~
160F, and then pa6ses by line 132 to cooler 133 where it is further
cooled to a tèmperature of 110F before introduction into the top of absorber
~ tower.
i In the absorber tower 69 the C02 content of the gas is reduced from ; ~ :
1.28~ to a residual concentration of 30 parts per million (ppm) by contact with
~ the diethanolamine solution, and the process gas leaving the top of absorber
J~ tower by line 73 has the romposition shown in Column 3 of Table II.
.~ The process gas stream is next conducted by line 73 to gas-to-gas
heat exchanger 74 where it is heated by the hot process gas fram methanation
unit 79 to a temperature of 228F and then i.s conducted by line 76 to gas-to-
gas heat exchanger 21 where it is further heated by contact with hot process
gas from high temperature shift converter 14 to a temperature of 600 F. It is
then introduced into methanation unit 79 where the small resldual amounts of ~ ~
C0 and C02 are converted to methane and water over a nickel methanation ~;
catalyst. The gas stream, now at a temperature of 604F and having the
composition shown in Column 1 of Table III is conducted by line 81 to waste
heat boiler 82 where it is cooled to a temperature of 300F and then to gas-to-
gas heat exchanger 74 where it is further cooled by indirect heat exchange with .
.~ the process gas coming from scrubbing tower 69 to a temperature of 192F. The
ga~ ~tr~am then flows to cooler 85 where it is further cooled to 100F. It
is then co~pressed to a pressure of 2200 psig and, after combining with recycle
''~, .,~

~,
:-' , .,

~3~5~
gas from line 96, is passed over a promoted iron ammonia synthesis catalyst
operating at an inlet temperature of 800 F and an outlet temperature of 960 F.
1 TABLE III
- Column 1 Column 2
Component Gas Composition Composition of Recycle
After Methanation and Purge Gas Vol %
Vol %
N2~Dry Basis) 24.85 20.22
H2 " " 74.52 60.63
- 10 CH4 ~ 0.32 8.15 ~- ~
A " " 0.31 7.74 `
3 ~~ 3.06
Total" " lO0.0~ 100.00 ;~

The process gas leaving the ammonia synthesis reactor containing ; ~`
. ,. : ,
15.17X NH3, 17.73N2, 53.16%H2, 6.79%Argon and 7.t5%cH4 is conducted to the
¦ ammonla recovery unit where the ~ulk of the ammonia product is ~emoved leaving
~ a recycle gas having the composit-lon shown ln Column 2 oE Table III. In order
.i . :
~' to maintain the combined concentration of methane plu8 Argon at a level of
16.1~ in the ammonia loop, a continuous purge gas stream is taken from the
20 recycle loop by line 98, such purge gas having the same composition as~the
recycle gas,.viz., that shown in Column 2 of Table III.
~¦ In order to illustrate the substantial reduction in process gas
1 losses without loss of overall thermal efficiency made possible through the `
;, process of the invention, the results obtained in the foregoing example are
l compared to the results obtained using identical quantities of process gas and
,, ,
3 process steam but employing an ammonia synthesis train operating according to'~ current normal practice. Such ammonia train differs from that employed in theinventlon in that intershift CO2 scrubbing is not employed; the process gas is
subjected to a high temperature and low temperature shift conversion followed -
by C02 scrubbing using an essentially isothermal hot potassium carbonate ~ ;
:, ;
-l scrubbing system after which the process gas is sub~ected to methanation and -~
~ s~nt to the ammonia synthesis loop. In Table IV the process gas composition -
i ~ust prior to methanation and following methanation, as well as the composi~ion

-33-
.,..

. i ' ,

~038~9~
of the purge gas, ls given for the ammonia train operated in accordance with theforegolng example and for an ammonia train operated according to such current
normal practice. ~` ;
TABLE IV
Component Gas Composltion Gas Composition Composition of
Prior to After Methanation Purge Gas
Methanation,Vol% Vol % Vol %
According According According
to Current to Currentto Current
Invention Practice Invention Practice Invention Practlce
N2(Dry Basi9)24.8224.10 24.85 24.65 20.22 20.09
H2 " " 74 5474.53 74.52 73.95 60.63 60.66
CH4 " " 0.30 .36 0.32 1.09 8.15 12.53
A " " 0.31 .31 0.31 .31 7.94 3.57 `~
C0 " " 0.003 0.1
2 (30ppm) tlOooppm) Nil Nll Nil Nil
C0 " " 0.002 0.60
~200ppm) (6000ppm) Nil Nil Nil Nil
NH3 ~~ 3.06 3.15
,j j :

As will be noted ~rom Table IV the combined level of C0 and C02 in
the gas ~ust prior to methanation is only 0.023% (230 ppm) in the example in
accordance with the invention while it is .41% (4100 ppm) in accordance with
current practice. After methanation the level of methane in the gas is only
-0.32% in the example according to the invention while it ia 1.09% in accordance ~ ;~
with current practice. The lesser amount of methane in the process gas
; reflects, of course, the lower hydrogen losses during the methanation step which
, ' '
characterizes operation in accordance with the invention. Even more importantly
" .
~, the Iower methane content of the process gas flowing to the ammonia synthesis ;~
loop means that there will be a much slower build-up of inerts in the ammonia
~ 30 synthesis loop and the amount of purging required to hold the level of inerts
; ~ at a tolerable level will be subetantlally reduced. In Table IY9 the total
concentration of inerts (i.e. methane plus argon) in the purge gas in the
i ~:
example accordlng to the invention and in the example according to current
practicé i~ the same, viz. 16.1% but in the example according to the invention

-34-
~3



... . . , , , . . ~ .

~s~ ~ .
~38~9~l ~
methane makes up only about half of this total, while in current practice
the methane makes up 77~ of the total.
The overall savings in process gas losses cmd the resulting
increased ammonia production from the same quantity of process feed gas can
be seen from Table V where the results from operating in accordance wlth the
foregoing example are compared to operation in accorclance with current ~ -
practice. -
TABLE V -

According to Current
Invention Practice

Process Gas ~eed, Methane, 2562 2562
pound mols per hour

; Proce~ Steam, total, pounds 184,650 184,650
per hour

Synthesis Gas Feed ~o Ammonia 11,035 10,7S2
~ Loop, pound mols per hour

;i Total Hydrogen Feed to Ammonia 8,224 7,952
Loop, pound mols per hour ;~

Percent Inerts in Feed Gas to 0.64% 1.4% `~
Ammonia Loop
Purge Gas Rate, pound mols per hour437 936

Hydrogen Lost in Purge Gas, 35.8 77.4 ; ~1 -
Tons per Day

Ammonla Lost in Purge Gas, 2.7 6.0
Tons per Day

Net Ammonia Production, Tons 1082 1000
- ~ per Day

., ~ .
Note that in these comparative examples, the amount of process gas
feed and the amount of process steam is identical in both cases. Due to the ~ -
30 lower losses of hydrogen in the methanation step, the amount of synthesis gas ;~ ~-
feed to the ammonia loop and the amount of hydrogen to the ammonia loop is
substantially increased in the operation according to the invention. Of
~ greater importance, the percent of inerts (methane plus argon~ in the feed gas
: t, to the ammonia loop is reduced to half the value of the percent inerts obtained
b~ ~peration in accordance with current practice, vi~. from 1.4Z inerts to

. r
'35~
'' , ~
. ~ .
-: ~
.
. - : . .... : - ~ . . . .:

1(~3~9~
0.64% inerts. The purge gas rate is correspondingly reduced to less than
half, viz. from 936 pound mols per hour to 437 pound mols per hour, while
both the hydrogen losses and ammonia losses are similarly correspondingly
reduced. Because of these reductions in the process gas losses, the net
production of ammonia is increased by 82 tons per day, a production increase
of 8.2% over current practice without any sacrifice of thermal efficiency. An
increase of 82 tons per day amounts to an increase of 25,000 tons per year
having a value of several million dollars per year.

~. , .
In Table VI the source of the savings in process gas losses when

operating in accordance with the foregoing example ls shown expressed in terms

of additional tons of ammonia production achieved.

TABLE VI

Source o~ IncreaseIncrease in Ammonia
Production Tons Per Day

Increase resulting from increased 8.8
hydrogen production in secondary
low temperature shift

Reduction in hydrogen loss resulting5.9
from methanation of carbon dioxide
., .
Reduction in hydrogen loss resulting26.3
~j from methanation of carbon monoxide

Reduction in purge gas quantity by 37.7
reduction of methane content of
synthesis gas
Reduction in ammonia losses in purge gas 3.3
Total increase in ammonia production82.0

While the invention has been illustrated and exemplified in terms
of ammonia synthesis trains e=ploying reforming processes to produce the
hydrogen-carbon monoxide mixtures which are then subjected to shift conversion,
it is to be understood that the invention is also applicable to ammonia synthesis
;~ trains where the hydrogen-carbon monoxide mixtures are produced by the partial
oxidation of hydrocarbons. Suitable hydrocarbon starting materials for
conversion to hydrogen-carbon monoxide mix~ures by partial oxidation include
. j
,~ ::~
~ gaseous hydrocarbons such as me~hane, ethane, propane, butane or mixtures of

,
-36-

::

~ ~:)38S9~
these; liquid hydrocarbons such as crude petroleum or petroleum fractions ~ ;
such as those in the naphtha range, the kerosene range, or the heavier
fractions) including the heavy residuals such as Bunker C; and solid hydro- ;
carbons such as coal, lignite or the like. Partial oxidation processes
to convert such hydrocarbon starting materials to mixtures of hydrogen and ;`
carbon monoxide unde~ super-atmospheric pressure are,of course,well known.
They generally involve reacting the hydrocarbon at elevated pressures with
oxygen and steam. Since many of such hydrocarbons contain sulphur means for
removal of the sulphur content of the process gas prior to contact with the
sulphur-sensitive low tempera~ure shift conversion catalysts should be provided. ; ;-
It is to be understood of course that the invention is not limited ~ ;
to the illustra~ive embodiments described herein and that other embodiments
not illustrated or specifically listed are included within the scope of the
appended clai~s.




.', ., ~,'',,",.''",



;l '`::
~ , . '~ ' '




. ' ,
`t '


- -37-


. ..
: i , , . . . . , .. .... .. . . : . . : :