Note: Descriptions are shown in the official language in which they were submitted.
9~2~
07-04Q3A
This invention relates to processes for separating at
least one gas from a gaseous feed mixture containing at
least one other gas by selective penmeation t~rough a
separation membrane
The need to separate at least one gas ~rom a gaseous
mi~ture is often encountered in modern so~iety. For
instance, the removal of contaminants in waste gas streams
may b~ xequired from an environmental standpoint, and
if the contaminants are useful, the removal and reco~ery
of the contaminants may be economicall~ desirable~ Moreover,
the recovery of one or more gases from a gaseous mîxture
may be a necessary procedure in chemical processing operations
Accordingly, many procedures have been d~veloped to ef~ect
gas separations such as selective condensation, adsorption-
desorption, absorption-desorption, and the like. One o~
the more recent proposals to eect gas separations is by
selective p~rmeation t~rough semi~permeable membranes, i~e~,
i separation m~mbranes.
According to current theories, gas separations effect~d
by separa~ion membranes may be by several mechanisms One
~` group of such mechanism~ include Knudsen flow, or dif~usion,
and the like which invol~e the passage of gases through
pores (i.e., continuous ~low channels for gas flow in
communication with both the feed and eæit suraces of the
membrane~ in the separation membrane. In another postulated
mechanism for gas separa~ions, the passage of ~ gas through
the membrane may be by interac~ion wi~h the material of
the membrane In order to effect the penmeation of a gas
through a separation mem~rane, a driving force must be
provided Generally, this driving force îs provîded by
'
.
L~29~26
-2 07-0403
maintaining a total pressure differential across the
thickness of the separation membrane~ Hence, the permeate
exit side of the separation membrane is often at a
substantially lower pressure than the feed side of the
separation membrane. The use of substantial total pressure
differentials is especially prevalent in connection with
g~s separation operations in which the permeation is by
interaction with the material of the separation membrane
in order to provide economically attractive fluxes of the
permeating gas per unit of available membrane surace area.
If, for instance, the permeating gas is to be
discharged to the environment or utilized at low pressure9
e.g., as a burner feed, the use of a substantial total
pressure differential across the separation membrane may
be wholly acceptable. However, it is oten desired to
employ the permeating gas in a chemical process operating
at superatmospheric pressure. For example, the gaseous
feed mixture to a separation membrane ma~ be an off-stream,
e g , a purge stream, from a superatmospheric synthesis
process using a cyclic reaction loop such as an ammonia or
methanol synthesis process~ At least one of the unreacted
reactants in the of~-s~ream may be recovered by permeation
through a separation membrane and returned to the super~
abmospheric synthesis process to enhance conversion yields
of the process. Thus, compression costs are incurred in
returning the permeating gas to the synthesis process.
These compression costs may off-set any savings which may
have been realized due ~o the recovery and returning to
~~ the synthesis process of the permeating gas.
Various methods for using a plurality of permeator
stages to effect the separation of a gas from a gaseous
feed mixture have baen suggested. For instance, United
States Patents Nos. 2,617,493 (November 11, 1952) and
3,713,271 (January 30, 1973) disclose cascade-type
permeator stages in which the permeating gas ~rom one
permea~or stage is passed to the feed side of a su~sequent
permeator stage. United States Patent No~ 3,339 7 341
(Septem~er 5, 1967~ discloses in connection with Figure 8
~129626
-3- 07-0403A
two permeator stages in seri~s in which the non-permeating
gas from the firs~ permeator stage is passed to the feed
side of th~ subsequent permeator stage; however~ the ratio
of total pressure at the feed side to total pressure at the
permeate exit side of the subsequent permeator stage is
disclosed to be lower than that ratio in the first permeator
stage. In West German published patent application
DT 26 52 432 (May 26, 1977) two permeator stages are
disclosed in which ~he non-permeating gas from the irst
permeator stage is passed to the feed side of the subsequent
permeator stage; however, the total pressure at the feed
side of each permeator stage is disclosed to be the same
and the total pressure at the permeate exit side of each
permeator stage is disclosed to be the same,
United States Patent No. 3,836,457 (September 17, 1974)
discloses a staged reverse osmosis system for purifying or
concentrating aqueous solution~ in which the concentrated
aqueous solution is passed ~o t:he feed side of a
subsequent reverse osmosis stage and the feed side of
the subsequent stage is operated at a higher total pressure
than a preceding stage; howe~er, no disclosure is provided
pertaining to the separation of gases.
Gardner, et al.~ in "Hollow Fiber Permeator for
Separating Gases", emical En~ineering Progress, October,
1977, pages 76 to 78, suggest that one application for
separation membranes is in treating an ammonia synthesis
purge stream to recover hydrogen. Gardner, et al., do not
disclose the use of permeator stages in series.
By this inven~ion processes are provided for separating
at least one gas from a gaseous feed mixture containing at
leas~ one other gas by selective permeation through a
separation membrane in which processes desirable ~mounts
of permeating gas can be obtained while requiring a reduced
amount of compression to provide the permeating gas at
advantageous elevated pressures. In accordance with the
processes of this invention a gaseous feed mîxture is
passed to at least two permeator stages in series. Each
of the permeator stages contain a separation membrane
1 1 2 '~
-4- 07-0403A
having a feed side and a permeate exit side and exhibiting
selectivity to the permeability of the at least one gas
as compared to the permeabilîty of the at least one other
gas. A total pressure differential is maintained across
t~ separation membrane to provide the driving force to
ef~ect the desired permeation of the at least one gas.
Between permeator stages, the non-permeating gas from the
feed side of the separation membrane of on~ permeator
staga is passed to the feed side of the separation membrane
of the next permeator stage~ The ratio of total pressure
on the feed side to total pressure on the permeate exit
side of the separation membrane for at least one penmeator
stage Chereinafter low total pressure ratio permeator
stage~ is less than the ratio of to~al pressure on the
lS feed side to total pressure on the permeate exit side of
the separation m~mbrane for at least one subsequent, i.e.,
downstream, permeator stage (hereinafter high total
pressure ratio permeator stage2.
In a highly advantageous use, permeator stages in
accordance with the invention are utilized to treat a purge
stream from an ammonia synthesis loop, Each permeation
stage contains a separation membrane which exhi~its selective
permeation of hydrogen as compared to the permeation o
inert contaminants in the purge stream. Hydrogen which
permeates through the separation membrane of at least on~
perme~tor stage can ~e rec~cled to the ammonia synthesis
reaction zone, The recover~ from ~he purge stream and
recycling of hydrogen to the ammonia synthesis reaction
can result in an en~anced conversion of ~ydrogen values
to ammonia~ This enhanced co~version of hydrogen values
can ~e achieved ~ven ~en process equipment design
limîta~ion do not permit an increase in the amount of
ammonia produced; however, often increased ammonia
production can ~e obtained,
T~e at least two permeator stages of this invention
provid~ significant advantages in t~at at least one low
total pressure ratio permeator ~tage separates the at
least one g~s ~rom the gaseous ~e~d mixtur~ whil~ ena~ling
the permeating gas from that ~tage to ~e at a desira~le
~ 1~9~20 -5- 07~0403A
total pressure which may require little, if any,
recompression for use in a chemical process, Thus, for
instance, in an ammonia synthesis process, an enhanced
CGnversion of hydrogen values provided by ~his invention
can involve little, if any, additional energy consumption
over similar ammonia processes which do not employ the
process of this invention, and in some instances the
energy consumption per unit ammonia produced is decreased.
The non~permeating gas from the at least one low total
pressure ratio permeator stage is passed to at least one
high total pressure ratio permeator stage in which
additional amounts of the at least one gas are separated.
Although the permeating gas from this permeator stage may
be at a lower total pressure than that of the permeating
gas from the at least one low total pressure ratio
permeator stage, the weight amount of permeating gas which
requires additional recompression is only a portion of the
permeating gases from all of the permeator stages. Thus,
less recompression is required than if all of the
permeating gases were at the lower total pressure. By the
use of the process of this invention, it is possibl~ to
enhance recovery of the at least one gas of the gaseous
mixture withou~ unduly increasing permeatîng gas
recompression costs. Moreover, the total available
separating membrane area for a given recovery of the at
least one gas is reduced using the processes of this
învention in comparison to t~e total available membrane
area required for ths given reco~ery of ~he at least one
gas if only low total pressure ratio permeators in
parallel flow relationship were employed~
~ ccording to current theor~, th~ rate at which a
moiet~ permeates through a separatîon membrane is
dependent in part on the driv7ng force for that moi~ty~
~ith respect ~o membrane separations in which the moiety
is gaseous and passes from a feed gas mixture to a
permeating gas on the exit side of the membrane, th~
driving force is the differential in fugacity for that
moiety. Generally, fugacities for ideal gases are
~ 129626 -6- 07-0403A
approximated by par~ial pressures and thus, conventionally,
in gas separations, the driving force is referred to in
terms of partial pressure differen~ials. The partial
pressure o~ a moi~y in a gas mixture can be defined as
the concentration of the moiety in the gas mixture on a
molecular basis times the total pressure of the gas mixture.
Often, the concentration of the moiety on a molecular basis
is approximated by the volume concentration of the moie~y.
In view of the effect of the concPntration of the moiety
in the gas and the to~al pressure of the gas on the partial
pressure, ~hese parameters can be varied jointly or
separately to provide suitable partial pressure differentials
across the membrane ~o provide desirable fluxes of the
moiety. For instance, with the moiety concentrations on
the feed side and on the permeate exit side and the total
pressure differential across ~he membrane remaining constantJ
but varying the total pressures on the feed and permeate
exit sides, a greater partial pressure differential of the
moiety is provided at lower total pressures on the feed side
and permeate exit side of the membrane,
Thu3, in accordance with this invention, the at least
one low total pressure ratio permeator stage can be operated
such that a suitable partial pressure differential for the
at least one gas is maintained across the separation
m~mbrane to provide, for ins~ance, a permeating gas
contaîning up to about 70 percent of the at least one gas
in the gaseous feed mîxture whereîn the permeating gas is
at a desirable total pressure for being used in a chemical
process without requiring undue rec~mpression. In certain
instances, it may be desirable to compress th~ gaseous feed
mîxture such that the permeating gas from this permeator
stage îs at a total pressure suitable for direct
reintroduction into the chemîcal process. In such instances,
the gaseous feed stream may of~en be compressed to at least
about 20 atmospheres above, say, about 25 to 100 atmospheres
above, the original pressure of the gaseous feed stream.
It is clear that th~ non-permeating gas from the low
total pressure ratio permeator stage will contain substantial
112g62'o
-7- 07-0403A
amounts of the at least one gas, for instance, at least
about 20 percent of the at least one gas in th~ gaseous
feed mixture. While additional amounts of the at least
one gas can often be recovered in the low total pressure
ratio permeator stage, e.g., by increasing the available
separation membrane area, it is preferred that this
permeator stage not be operated to maximize its recovery
of the at least one gas. Rather, this permeator stage is
preferably operated predominantly on a flux-limiting basis.
In a flux-limiting basis operation, the separation is
conducted u~der conditions such that when the flux of
the at least one gas through the membrane significantly
decreases, the separation operation is terminated, e.g.,
by p~ssing thP non-permeating gas rom the permeator~
Flux-limiting basis operations are in contrast to unwanted
permeate~limiting basis operations, In unwanted permeate-
limiting basis operations, the separation is continued to
provide a suitable recovery of a high proportion of the
moiety from the feed mixture without undue permeation of
the undesired moieties in the fl_ed mix~ure. Generally1
in any commlercially practical mlembrane separation operation,
both flux-limiting basis and unwanted permate-limiting
basis considerations will be involved, Often, in a
predominantly flux-limiting mode of operation, it is
desired that the percent of the difference in partial
pressures of the at least one gas (A~ between the gaseous
eed mixture CppA feed~ and the non-permeating gas
CppA non~permeating) divided by the difference ~etween
the partial pressure of the at least one gas in the gaseous
feed mixture and the minimum partial pressure of the at
least one gas on the permeate Pxit side of the membrane
CppA permeate min.~ is up to about 90, say, about 20 or
30 to 90, often a~out 30 to 85. On the other hand, in a
predominantly unwanted permeate-limiting basis mode of
operation, this relatîonship will often be at least about
85 or ~0 percent.
As stated above, the low total pressure ratio
permeator stage is preferabl~ opera~ed on a predominantly
1129~6
-8- 07-0403A
flux-limiting basis in order to provide a permeating gas
at a desirable total pressure, For a given total pressure
differential across the separation membrane and a given
separation membrane, a high purge stream flow rate per unit
S of available membrane surface area can be employed and a
greater amount of the at least one gas permeates the
membrane per unit area per unit time than if the permeator
stage were operated on an u~wanted permeate-limiting basis.
Generally, sufficient membrane area is provided in the
low total pressure ratio permeator stages to permeate at
least about 20, preferably about 30 to 70, percent of the
at least one gas in the gaseous feed mixture,
Since the low total pressure ratio permeator stages
are preferably flux limited, particularly desirable
separation membranes exhibit high permeabilities for the
permeation of the at least one gas, but need not exhibit
as high a selectiv.i~y to the pe:nmeability of ~he at least
one gas as compared to ~he permeability of the at least
one other gas in the gaseous mixture as the selectivity
required of a membrane in a predominantly unwanted permeate-
limited mode of operation or if the separation were
conducted in a single permeator stage to provide the same
overall recovery of the at least one gas,
The non-penmeating gas from the at least one low
total pressure ratio permeator s~age is passed to the feed
side of at least one high total pressure ratio permeator
stage to recover additional amounts of the at least one gas.
The amount of the at least one gas in ~he permeating gas
from this permeator s~age is frequentl~ at least about 10,
say, at least about 15 percent of the amount of the at
least one gas in the gaseous feed mixture. The amount of
the at least one gas in the total permeating gas fro~ all
of the permeator stages is preferably at least about 50,
e,g., at least about 60, say, a~out 6Q to 95, percent of
the at least one gas in the gaseous feed mi~ture.
The at least one high total pressure ratio permeator
stage can be operated on a predomînantly flux-limited
basis or a predominantl~ unwanted permeate~limited basis.
~.~29626
~9- 07-0403A
The gas fed to the high total pressure ratio permeator
stage can be at any suitable total pressure, For instance,
the non-pe~meating gas from the low total pressure ra~io
permeator stage can be compressed or decompressed, or can
remain at substantially the same pressure depending upon
the desired total pressure differential across the
separation mem~rane, the total pressure of the permeating
gas, and the like. Often, due to strengths obtainable in
some suitable separation membranes, the total pressure of
the gas fed to the high total pressure ratio permeator
stage is decompressed to enable achieving a desirable
total pressure differential across the membrane.
The ratio of the total pressure on the feed side to
the total pressure on the permeate exit side of the at
least one low total pressure ratio permeator stage is less
than that ratio for the at least one high total pressure
ratio permeator sta~e. Often, the total pressure ra~io
of at Least one low total pressure ratio permeator stage
is at least about 10 or 15, say, about 15 to 9~, preferably,
about 20 to 95, percent less th,an the total pressure ratio
of at least one high total pressure ratio permeator stage.
Generally the total pressure drop across at least one high
total pressure ratio permeator stage is wQthin about 10 to
500, say, about 15 to 250, percent of the total pressure
drop across at least one low total pressure ratio permeator
stage. In one aspect of this invention, the total pressure
on the parmeate exit side of ~he higher total pressure ratio
permeator s~age is at a lower total pressure than the total
pressure on the permeate exit side of the lower total
pressure ratio permeator stage,
An~ suitable number of permeator stages may be employed
so long as at least one low total pressure ratio permeator
stage and at least one high total pressure ratio permeator
stage are provided. Each permeator stage may be comprised
of one or more separate permeators wherein plural permeators
are arranged in substantially parallel flow relationships.
Preferably, the first permeator stage is a low total
press~re ratio permeator stage, Often, the last permeator
11~96~
-10- 07-04Q3A
stage is a high total pressure ratio permeator stage.
Most frequently, two permeator stages are u~ilized, however,
in some instances three or more permeator stages mar be
desirable. Generally, little ben~fit is achieved in the
use of permeator stages above about ive. Preferably, if
any permeator stage is operated on a predominantly unwanted
permeate-limiting basis, that permeator stage is the last
permeator stage.
The effective membrane surface area ci.e., the membrane
area a~ailable to effect separation) for each permeator
stage should be sufficient to allow a desired amount o~
the at least one gas to permeate. The amount of effective
membrane surface area to be employed is influenced ~y, for
instance, the permeation rate o~ the at least one gas through
the membrane under the separation conditions, i.e.,
temperature, absolute pressure, total pressure differential
across the membrane, and partial pressure dif~eren~ials of
the at least one gas across t~e membrane, Advantageous
total pressure differentials across separation membranes
are at least abouk 10, say, at least about 2~, atmosp~eres 2
and ma~ be up ~o 100 or 200 atmospheres or more. However,
the pressure differential shou].d not be so great as to
~nduly stress the membranes suc~ t~at it ruptures or is
prone to easily rupturing.
2S A permeator containing the separation membrane may
be of any suitable design for gas separations, e.g~, plate
and frame, or ha~ing spiral wound film membranes> tu~ular
membranes, hollow fi~er membranes, or t~ like. Preferably,
the permeator compxises hollow fi~er membranes due to the
high membrane surface area per unit volume w~ich can b~
obtained, ~en the membranes are in tubular or hollow fiber
form, a plurality of the membranes can be substantially
parallelly arranged in bundle form and the gaseous feed
mixture can be contacted with either the ou~side (shell
side) or the inside (bore side) of the membranes.
Preferably, the gaseous feed mixture is contacted with the
shell side of the membranes since passage of the gaseous
feed mixture through the bore side of the membranes may
9~26
~ 07-0403A
involve substantially greater pressure losses. With shell
side feed, the shell side effluent from the permeator can
often be at less than about 1 or 5, often within less than
about 0.5, atmospheres below the pressure of the gaseous
feed mixture fed to the permeator and thus be at an
advantageous pressure for subsequent processing or energy
recovery, e.g., by the use of turbines. Since the
concentration of ~he at lea~t one gas on the feed side o~
the membrane is continually diminishing as the at least one
gas permeates to the permeate exit side of the membrane which
has increasing concentration of the at least one gas, the
partial pressure differential of the at least one gas across
the membrane is continually changing~ Therefore, flow
patterns in the permeator can be utilized to provide desirable
recoveries o the at least one gas from the gaseous feed
mixture. For instance, the flows of the gaseous feed mixture
and the permeating gas can be concurrent or countercurrent~
With bundles of hollow fiber and tubular mem~ranes, the
shell side feed can be radîal, i.e , the feed stream
transversely flows past the membrane either to the insid~
or, usually the outside of the bundle, or the flow can be
axial, i.e., the feed stream disperses within the bundle
and generally flows in the direction in which the hollo~
fibers or tubular m~mbranes are oriented
Any suitable material ma~ be emplo~ed for t~e
separation m~m~rane as is well-known in the art Typical
membrane materials include organic polymers or organic
polymer mixed with inorganics, e,g~, fillers, reinforcements,
and the like. Metallic and metal-containing membranes may
also be used.
The ammonia synthesis process is descri~ed in more
detail in order that this aspect invention can be full~
appreciated~ Ammonia is synthesized b~ the catalytic
reac~ion of hydrogen and nitrogen~ The hydrogen feedstock
for the ammonia synthesis is generally obtained from prîmary
reforming of hydrocarbon, etg,, natural gas The e~fluent
from the primary reforming thus contains impuriti~s such
as methane, carbon oxides, i~e , car~on dioxide and car~on
1 .1.296~5 -12- 07-0403A
monoxide, wa~er and the like. Current practîce provides
for the removal of impurities from the reformer effluent
which may be harmful to the ammonia synthesis catalyst such
as the carbon oxides, sulur compounds and the like; however,
impurities such as methane are generally not completely
removed from the reformer effluent since they are not
directly harmful to the ammonia synthesis reaction and are
expensive to remove. The nitrogen feedstock is usually
obtained rom air with the removal of oxyg~n, e.g,, b~
combustion wlth fuel to produce water or carbon dioxide and
water, followed by removal of the water and carbon di~xide,
if present, or by liquifaction. The resultant nîtrogen
stream contains minor amounts of ~mpurities such as argon
which are present in small amounts in air, Since they are
not directly harmful to the ammonia synthesis reaction,
these impurities are not generalLy removed from the nitrogen
feedstock due to economic considerations, Thus, even though
the predominant components of the synthesis eed gas are
hydrogen and nitrogen, at least on~ of methane and argon are
present as contaminants in the synthesis feed gas, Me~hane
is often present in amounts of up to about 5, e,g,, about
0.1 to 3, volume percent, and argon is often present in
amounts of up to about 0,5, e,g., about 0,1 to 0,5, mos~
often about 0,3, volume per~ent based on the synthesis feed
gas. Othe~ contaminan~s which may be present include water
and helium.
The ratio of hydrogen to nitrogen which is present in
the synthesis feed gas is preferabl~ such that the mole ratio
of hydrogen to nitrogen of the reaction gas introdueed into
the ammonia synthesis reaction zone is su~stantially constant
to prevent a build-up of either ~ydrogen or nitrogen in the
ammonia synthesis loop. However, the mole ratio of hydrogen
to nitrogen in the reaction gas may ~e greater or less than
the stoichiometric ratio such that the excess of hydrogen
or nitrogen over that required for the reaction to ammonia
on a stoichiometric basis shifts the equilibrium in favor
of ammonia production. In such situations, the mole ratio
of hydrogen to nitrogen may ~e from about 2 or 2~5:1 to
9~26
- 13 -
about 3.5 or 4:1. Higher or lower mole ratios could be
employed; however, since a purge stream must be removed
from the synthesis loop to prevent undue build-up of
contaminants, considerable increases in the loss of valuable
nitrogen or hydrogen would be incurred. The processes of
this invention do minimize the increases in loss of hydrogen
through the purge stream when the reaction gas has a greater
than 3:1 mole ratio of hydrogen to nitrogen because of the
recovery and recycling of hydrogen from the purge stream.
Generally, the mole ratio of hydrogen to nitrogen in the
reaction gas is about 2.8:1 to 3.5:1, say, about 2.9:1 to
3.3:1. Frequently, the mole ratio of hydrogen to nitrogen
in the reaction gas introduced into the ammonia synthesis
reaction zone is substantially that mole ratio required for
the reaction of hydrogen and nitrogen on a stoichiometric
basis, e.g., about 2~95:1 to 3.05:1. Generally, nitrogen
does not permeate the membrane to a significan-t extent, and
the permeating gas contains little, if any, nitrogen. How-
ever, any nitrogen which is recovered and recycled in the
permeating gas represents a savings with respect to the nitro-
gen Eeedstock demands. The mole ratio of hydrogen to nitro-
gen in the synthesis feed yas is thus usua:Lly slightly less
than the mole ratio of hydrogen to nitrogen in the reaction
gas such that desirable hydrogen to nitrogen ratios are
provided when combined with the permeating gas which is re-
covered from the purge stream. In typical ammonia plants in
accordance with this invention, the mole ratio of hydrogen
~~ to nitrogen in the synthesis feed gas may be about 2.7:1 to
3.2:1, say, about 2.3:1 to 3.0:1.
The reaction between hydrogen and nitrogen to produce
ammonia is exothermic and is an equilibrium reaction. The
ammonia synthesis may be conducted using any suitable pro-
cedure such as the Haber-Bosch, modified Haber-Bosch, Fauser
and Mont Cenis systems. See, the Encyclopedia of Chemical
Technology, Second Edition, Volume 2, pages 258, et seq., for
1~2962~`
- 13a -
various of the processes for synthesizing ammonia from
hydrogen and nitrogen. In general, these processes employ
super-atmospheric ammonia synthesis pressures of at least
about
~I
6 ~ ~
-14- 07-0403A
100 atmospheres absolute and promoted iron synthesis
catalysts. The ammonia synthesis reaction zone is generally
cooled to maintain reaction temperatures of about 150 or
200 to 600C. The use of high synthesis pressures shifts
the equilibrium in favor of the formation o ammonia.
Although some ammonia synthesis pressures which have been
employed are as high as 500 or more atmospheres absolute,
most present day ammonia plants utilize synthesis pressures
of about 100 to 300 or 350 atmospheres absolute, especially
about 125 to 275 atmospheres absolute, Typically, the
ammonia synthesis feed gas is compressed in at least two
stages in order to facilitate achieving synthesîs pressures.
Generally, the pressure of the feed gas prior to at least
one compression stage is within at least about 100, say,
within about 10 or 20, atmospheres below the synthesis
pressure. The lowest pressure in the ammonia synthesis
loop is preferably within about 5 or lO atmospheres below
the synthesis pressure. A recycle compressor is generally
emplo~ed to circulate the gases in the synthesis loop and
to main~ain the desired synthesis pressure in the ammonia
synthesis reaction zone.
The conversion to ammonia based on hydrogen entering
the ammonia syn~hesis reaction zon~ is often a~out 5 to 30,
e~g., about 8 to 20 percent~ In many commercial plants, the
ammonia concentration of the reaction effluent exiting the
ammonia synthesis reaction zone is about lO to 25, e.g.,
about 10 to 15 or 20, volume percent~ Thus the reaction
effluent from the ammonia synthesis reaction zone contains
substantial amounts of hydrogen and nitrogen. Accordingly,
ammonia is condensed from the reaction effluent, and the
reaction effluent containing the valuable hydrogen is
recycled in an ammonia synthesis loop to the ammonia
synthesis reaction zone to provide an attracti~e convQrsion
of hydrogen in the feed to ammonia. Frequently, the reactor
feed gas fed to the ammonia synthesis reaction zone contains
about 0.5 to 5, say, about 1 to 4, volume percent ammonia
and less than about 25 volume percent inert contaminants,
say, about 4 to 15 volume percent inert contaminants
:1~12962~
-15- 07-0403A
T~us, th~ reactor feed gas may comprise about 2 to lS
volume percent methane, about 2 to lO volume percent
argon, and helium, if present in the reformer feed, e.g.,
in an amount of about 0.1 to 5 volume percent.
Ammonia in the reaction effluent from the ammonia
synt~esis reaction zone is removed from the synthesis loop.
A preferred method for removing the ammonia is by chilling
the ammonia-containing reaction effluent to coalesce ammonia
which can be removed as a liquid product After removal of
the ammonia the gas în the synthesis loop still may contain
ammonia, e.g,, up to about S volume percent ammonia. The
coalescing of ammonia from the gas in the ammonia synthesis
loop is preferably conducted subsequent to the recycle
compression. Two or more ammonia coalescers may be employed
in the synthesis loop to enhance ammonia recovery.
The compressed synthesis feed gas may be introduced
into the ammonia synthesis loop at any suitable location,
e.g,, before or after the recycle compressor, and before or
after the ammonia removal, In many instances, however, it
is preerred to introduce the compressed synthesis feed gas
into the ammonia synthesis loop prior to coalescîng ammonia
since the coalescing can remove water vapor and thus ensure
that the reaction gas fed to the ammonia synthesis reaction
zone h~s a low oxygen-containing compound content to prevent
catalyst poisoning.
Difficulties occur in that the inert contaminants such
as methane, argon, etc., in t~e hydrogen and nitrogen
feedstocks do not take part in the ammonia synthesis
reaction and must be removed from the ammonia synthesis
loop in an amount suficient ~o prevent an undue build-up
of these inert contaminants in the ammonia synthesis loop.
Conveniently, the removal of these inert contaminants is
effected by removing a purge stream Erom the ammonia
synthesis loop. The purge stream will contain the same
concentration of hydrogen and nitrogen as the recycling
reaction effluent. Hence~ recovery o the valuable hydrogen
from the purge stream for return to the ammonia synthesis
catalyst zone may be highly desirable, Frequently, the
~.29~
-16~ 07~0403A
reactor feed gas contains less than about 25, say, about 4
to 15, volume percent inert contaminants. The purge stream
often comprises up to about 3, say, about 0.5 to 2.5, volume
percent of the gases in the synthesis loop at the point from
which the purge is taken. The purge stream may, o course,
be a greater portion of the gases in the synthesis loop;
however, such large purge amounts result in increases in the
weight amounts of nitrogen and, possibly, hydrogen exhausted
from the ammonia synthesis system, The volume of the purge
stream is usually sufficient to maintain the concentrations
of methane and argon substantially constant.
It is generally preferred to remove the purge stream
from the gases in the ammonia synthesis loop upstream of the
introduction of the compressed synthesis ~eed gas to prevent
purging the fresh hydrogen and nitrogen feed. The purge
stream may be removed from the synthesis loop upstream of the
ammonia removal, or the purge stream may be removed ~rom the
synthesis loop downstream of the ammonia removal from the
synthesis loop. Usually the gases in the ammonia synthesis
loop downstream from the ammonia removal contain reduced,
but still significant, amounts of ammonia.
In the case in which the purge stream is removed rom
the synthesis loop upstream of the ammonia removal, the
ammonia concentration in the purge stream is often at least
about 5 volume percent, say, up to about 30, e.g., about 8
to 25, or e~en 10 to 15 or 20, volume percent. Convenlently,
the purge stream is chilled to coalesce ammonia, and the
separated liquid ammonia can provide additional ammonia
product The purge stre~m still contains sîgnifîcant
amounts of ammonia, e.g., often at least about 0 5 or 1
volume psrcent ammonia. This procedure is par~icularly
desirable w~en modifying existing ammonia synthesis plants
to produce ammonia in accordance with this in~ention since
existing ammonia synthesis plants generally emplo~ an ammonia
coalescer to remove ammonia from the purge stream. The
amount of ammonia in the purge stream ma~ be further reduced
b~ scrubbing with water or by diffusion of the ammonia
through the separation membrane, ~hus, the ammonia
.L 1 ~ 3 ~ ~ S
-17- 07-0403A
concentration of the non-permeating gas from the last
permeator stage may be sufficîently low that it is suitable
for use as, e.g., fuel,or can be vented to the environment,
especially after recovering energy provided by the higher
pressure of the purge stream. On the other hand, the purge
stre~m may ~e passed to the permeator stages without removal
of ammonia, or ammonia can be removed from the non-permeating
gas between permeator stages, e.g., by chilling and
coalescing and/or b~ water scrub~ing.
In the case in which the purge stream is removed from
the synthesis loop downstream of the ammonia r~moval, the
ammonia concentration in the purge stream is often at least
about 0 5 up to about 5 volume percent. Tn view of the low
ammonia concentration in the purge stream, removal of ammonia
from the purge stream prior to contacting the separation
membrane sometimes is not done. Additional ammonia is
recovered from the purge stream by permeation through the
separation membrane, and the non-permeating gas from the
last permeator stage may be suitable for use as, e,g., fuel
or can be vented to the atmosphere, especially after
recovering energy provided by the high pressure of the purge
stseam,
The purge stre~m may, if necessary, be subjected to
heat exchange to pro~ide suitable temperatures for efIecting
hydrogen separation by the use of separation m~mbranes
Often, the purge stream to ~e contacted with t~e separation
membran~ o~ a permeator stage is at least a~out 10C, say?
about 15 to 50C, preferably, about 25 to 40C, Higher
temperatures may be employed dPpending upon the physical
stability and thè selectivity of separation of t~e mem~rane
at the higher temperatures.
The purge stream is contacted with a separation
membrane which exhibits selectivit~ to the permPat;on o~
hydrogen as compared to the permeation of each of m~t~ane
and argon. In viPw of the generally substantîally lower
volume concen~rations of methane and argon in the purge
stream as compared to the volume concentration o~ hydrogen
in the purge stream, suitable separation mem~ranes need
-
11~96~
-18- 07-0403A
not exhibit high selectivit~ of separation of hydrogen from
each of methane and argon in order to provide an enhanced
ammonia synthesis process~ Generally, the selectivity of
separation of a membrane is described in terms of the ratio
of the permeability of the fast permeating gas Chydrogen) to
the permeability of the slow permeating gas (methan~ or
argon~ whereîn the permeability of the gas through the
mem~rane can be defined as t~e volume of gas, standard
temperature and pressure, which passes through a membrane
per square centimeter of surface area, per second, for a
partial pressure drop of 1 centimeter of mercur~ across the
membrane, This ratio is referred to as a separation factor
for the membrane, For sake of uniformit~, the permeabîlities
and separation factors mentioned hRrein are determined a~
a~out 25C and a pressure drop o a~out 3.4 atmospheres
across the membrane with the feed side of the membrane ~eing
about 3.4 atmospheres absolute unless otherwise indicated.
Often, the separation factor of the m~mbrane for the
separation o~ ~lydrogen from methane is at least about l
Separation ~actors ~or hydrogen over methane of 100 or
greater ma~ be provided by certain membranes; however, lit~le
advantage ma~ be ohtained us~ng such highl~ selecti~e
mem~ranes, Often the membrane may be sPlected on its
ability to quirkly permeate hydrogen rather than on its
selectivity of separation. Consequently, membranes
exhibiting a separation factor for hydrogen over methane
of about 10 to 80 are adequate. Clearly, the higher the
permeability of hydrogen through a membrane, the less
available membrane surface area which is required to pass
a desired amount of hydrogen through the membrane,
Particularly desirable membranes exhibit hydrogen
permeabilities of at least abou~ 1 x 10-6, preferably at
least about 20 x 10-6, cubic centimeters of hydrogen per
square centimeters of m~mbrane surface area per second at
a partial pressure drop of 1 centimeter of mercury across
the membrane.
1129~6
-19- 07-0403A
The volume ratio of the permeating to non-permeating
gases from each of the permeator stages as well as the
composition of each of the permeating and non-permeating
gases which may be employed in accordance with the method
o~ this invention can be ~aried over a wide range. By way
of illustration, Table I provides typical approximate
concentrations of the significant components in the gases
passed to a low total pressure ratio permeation stage and
to a high total pressure ratio permeator stage and those in
the permeating and non-permeating gases from each stage.
The permeating gas from each of the permeator stages
contains valuable hydrogen and can be recycled such that
the hydrogen can be utilized in the ammonia synthesis. In
accordanc.e with the processes of this invention the total
pressure under which the permeating gas exits each permeator
stage is taken advantage of, for instance, by returning the
permeating gas to the synthesis feed gas at a point where
the permeating gas is at substantially the same pressure as
the s~nthesis gas. Thus recompression costs are minimized.
The stream into which a permeating gas is introduced
can be selected partially on the basis of the operating
pressure differentials across t:he separation membrane which
can be employed. Since the compression of the synthesis
feed gas is condu~ted in severaLl stages, or steps, some
limitation exists as to the pressure differentials across
the membrane which are available in a given ammonia synthesis
system, especially in ammonia synthesis systems which are
retrofitted with separation membrane hydrogen reco~ery
sy~tems to enable conducting the ammonia synthesis processPs
of this ~nvention. The processes of ~his invention,
however, are sufficiently flexible, since ~he separation
of h~drogen is conducted in a plurality of permeator stages,
that a permeating gas stream having a desirable total
pressure can be provided. In general, the pressure
differential for a given ammonia synthesis system in
a~cordance with this invention is selected to pro~ide the
largest operating total pressure differential across the
separation membrane (within the range of suîtable operating
9 6 ,~ ~ -20- 07-0403A
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1 ~9~ 21- 07-0403A
pressure differentials for a given separation membrane)
which provides a permeating gas at a suitable pressure for
introduction into a synthesis feed gas stream. Usually,
the permeating gas is at a slightly higher pressure than
the gas stream into which it is introduced, e.g., abou~ 0.1
to 5 atmosph~res higher. A reduction in the total pressure
of the permeating gas or, preferably, the gas passing the
permeator stage should only be employed when no suitable
gas stream is available which is at a pressure which permits
a suitable total pressure differential across the membrane
to be achieved, e.g., the available synthesis feed gas
streams are at pressures too high to permit a desired
hydrogen flux through t~e separation membrane or at
pressures so low that the separation membrane can not
physically withstand the pressure differential. Alternatively,
a desirable total pressure differential may be provided
across the separation membrane and the permeating gas
compressed to suitable pressures for introduction into>
e.g., the synthesis feed gas stream or synthesis loop, or
the puxge stream may be compressed prior to contacting the
separation membrane to provide a desired pressure diferential
across the separation membrane and a permeating gas at a
suitable pressure for reintroduction into the ammonia
synthesis system.
The non-pe~meating gas from the last permeator stage
may be utilized in any suitable manner, e.g., used as fuel.
Since the non-permeating gas is at high pressure, significant
energy can be recovered from this gas by, for instance, the
use of a ~urbine, and the like.
The invention will be further described with reference
to the drawings in which:
FIGURE 1 is a simplified schematic flow diagram of
an ammonia synthesis plant having two permeator stages in
accordance with ~his invention wherein the permeating gas
rom the first permeator stage is combined with the
synthesis eed gas between compression stages for recycling
to the ammonia synthesis reaction zone and the permeating
11~9626 -22- 07-Q403A
gas from the second permeator stage is com~ined with the
synthesis feed gas prior to its compression to the super~
atmospheric synthesis pressure.
FIGURE 2 is a simplified schematic flow diagram of an
ammonia synthesis plant having two permeator stages and is
similar to the plant depicted in Figure 1 except that the
permeating gas from the second permeator stage is compressed
and combined with the permeating gas to ~e recycled to the
ammonia synthesis reaction zone~
FIGURE 3 is a simplified schematic flow diagram of an
ammonia synthesis plant having tWQ perme~tor stages in
accordance with this invention wherein ammonia is removed
from the ammonia synthesis loop prior to the removal o~ the
purge stream, The permeating gas from the first permeator
stage contains ammonia and is combined with dry synthesîs
feed gas between compressor stages, The non-permeating gas
from the first permeator stage is passed through a scru~ber
to remove ammonia and then to the second permeator stage
rom which permeator s~age the permeating gas is combined
with the synthesis feed gas pri.or to its compression to the
superatmospheric synthesis pre~,sure,
FIGU~E 4 is a simpli~ied schematic flow diagram of an
ammonia synthesis plant having two permeator stages and is
similar to the plan~ depicted in Figur~ 3 except that ~he
purge stream is compressed prior to ~eing passed to the
first permeator stage and the permeating gas from th~ ~irst
permeator stage is directly introduced into the ammonîa
synthesis loop,
FIGURE 5 is a schematic cross-~ection of a holLow
fiber mem~rane-containing permeator which may ~e employed
in an ammonia synthesis plant in accordance with this
invention.
11'~9626
-23- 07-0403A
In Figures 1 and 2 and 3 and 4, like reference numerals
indicate like features.
With reference to Figure 1, a synthesis feed gas
comprising hydrogen and nitrogen in approximately a 3:1
mole ratio is introduced via line 10 into the ammonia
synthesis system, The synthesis feed gas is compressed
in several stages to superatmospherîc ammonia synthesis
pressures. As depicted, compressor 12 partiall~ elevates
the pressure of the synthesis feed gas toward ~e super~
atmospheric synthesis pressure. The partially compressed
feed gas is chilled in coalescer 14 to condense and remove
(:~ia line 15~ water vapor from the synt~esis feed gas. The
thusly dried synthesis feed gas is passed to compressor 16
where it is compressed to above the superatmospheric
pressure in the ammonia synthesis loop, It is ~o ~e
realized that one or both of compressors 12 and 16 may ~e
comprîsed of two or more compressor stages,
The effluent from compressor 16 is transported via
conduit 18 to the ammonia synthesis loop. The conventional
synthesis loop is depicted in t:hat the s~n~hesis feed ga~
passes through rec~cle compressor 20 and ammonia coalescer
22 from which product a~onia is withdrawn via line 23 before
entering ammonia synthesis reaction zone 24 for con~ersion
to ammonia. Positioning coalescer 22 prior to the ammonia
synthesis reaction zone 24 insures that any water vapor
which may be present in the reactor feed gas is reduced to
provide less than about 10 ppmv total oxygen-containing
compounds in the reactor feed gas, The reaction ef~luent
from the ammonia synthesis reactîon zone is cooled in heat
exchanger 26 to a t~mperature in t~e range of a~out 0 to
100C, The heat ~ransfer medium in heat exc~anger 26 ma~
~e the reactor feed gas from ammonia coalescer 22 which in
turn is heated to a suitable temperature for introduction
into the ammonia synthesis reaction zone. T~e efluent
from heat exchanger 26 is recycled via line 28 Cammonia
synthesis loop~ to the recycle compressor 20,
A purge stream is withdrawn rom line 28 via line 3Q,
The volume of the purge s~ream is sufficient to main~ain an
11~9~26
-24- 07-0403A
acceptable level of inert contaminants in the ammonia
synthesis loop and ammonia synthesis reaction zone. As
depicted, the purge stream is removed prior to the
condensation of the ammonia product from the gases in the
ammonia synthesis loop. Thus, the purge stream contains
substantial quantities of ammonia. As depicted in Figure 1
ammonia is removed from the purga stream by passing the purge
stream throug~ chiller-coalescer 32 which removes ~mmonia by
condensation followed by water scrubber 33 which absorbs
ammonia and generally provides a gas containing less than
about 0.1 volume percent ammonia~ The purge stream may,
if necessary, be subjected to heat exchange to provide a
temperature of, say, about 25 ~o 40C. The purge stream
is passed via line 34 to first permeator 36.
A schematic cross-section of an axially, shell-side
fed permeator such as may be employed in the system of
Figure 1 i9 provided in Figure 5. With reference to
Figure 5, within casing 100 is positioned a plurality of
hollow fiber membranes which are arranged in bundle
generally designated by the numeral 102, One end of the
bundle is emBedded in header 104 such tha~ the bores of
the hollow fibers communicate through the header. The
header is positioned in casing 100 such that essentially
the only fluid communication t~rough the header is through
the bores of the hollow fi~ers. The opposite ends of the
hollow fibers are sealed in end seal 106, Th~ purge stream
enters the casing through feed port 108, disperses within
bundle 102 and passes to shell exît port 110 positioned at
the opposite end of the casing~ ~ydrogen permeates to the
bores of the hollow fibers, and passes via t~e bores through
header 104, The permeating gas exits casing 100 through
permeate exit port 112. While Figure 5 depicts a hollow
fiber membrane~containing permeator in which onl~ one end
of the holl~w fibers is open, it is apparent that both ends
35 of the hollow fi~ers can be open.
~ ith reference to Figure 1, a first permeating gas~
i,e,, a hydrogen-rich stream, exits permeator 36 via line
38~ Th2 pressure drop across the membrane is such that the
ll~9S~
-25- 07-0403A
permeating gas is at a pressure substantially the same as
the pressure of the synthesis feed gas exiting compressor
12, and the first permeating gas is combined with the
synthesis feed gas exiting compressor 12 in order to be
recycled to ammonia synthesis reaction zone 24, The first
permeating gas is introduced into the synthesis feed gas
upstream of coalescer 14 such that water vapor which is
introduced into the purge stream in scrubber 33 and permeated
through the separation mem~rane, can be removed
The non-permeating gas is withdrawn from the feed side
of first permeator 36 and is passed via line 40 to permeator
42. The non-permeating gas contains hydrogen as well as
nitrogen, methane and argon. A second permeating gas exits
second permeator 42 via line 44. The second permeating gas
is at a pressure substantially the same as the pressure of
the synthesis feed gas entering compressor 12, and the second
permeating gas is combined with the synthesis eed ~as
entering compressor 12 to be recycled to ammonia synthesis
reaction zone 24. The non-permeating gas from t~e second
permeator exits via line 46 and can be treated in an
additional permeator (not depicted), exhausted to the
environment, or used, for instance, as a fuel.
The ammonia synthesis syst~m of Figure 2 îs substantially
the same as the system depicted in Figure 1 except t~at the
second permeating gas from second permeator 42 is compressed
in compressor 48 to a total pressure slightly above the
pressure of the first permeatîng gas in line 38. The
compressed second penmeating gas is passed via line 50 to
line 38 whereat it is combined wit~ the first permeating gas
~eîng recycled to ammonia synthes~s reaction zone 24~ This
method may find applicatîon when retrofitting separation
membrane ~ydrogen recovery systems in exîsting ammonia plants
in order to utilize a process of this invention~ For
instance, îf încreased ammonia production is desired in an
ammonia plant, ~ut compressor 12 is at its maximum capacity,
thîs ~ottleneck can be obvîated ~y utilizing a compressor
to increase the pressure of the second permeatin~ gas such
that it can be introduced into the synt~esis feed gas
1~.2962~
-26- 07-0403~
without increasing the load through compressor 12, Also,
the difference between the pressure on the feed side of
second permeator 42 and the feed side of compressor 12, in
some ammonia plants, may bP too great to be withstood by a
membrane which may be employed. Accordingly, a lesser
pressure drop can be utilized across the membrane in the
second permeator while maintaining the feed side o~ the
second permeator at elevated pressures~ In view of the use
of high pressures on the feed side of the second permeator,
little compression is required to elevate the pressure of the
second permeating gas for introduction into the ammonia
synthesis system.
In the ammonia synthesis system depicted in Figure 3
the purge stream is withdrawn from the ammon~a synthesis
loop downstream of the ammonia removal and upstream of the
introduction of the fresh synthesis ~eed gas into the
synthesis loop,
A synthesis feed gas comprising hydrogen and nitrogen
is introduced via line 200 into the ammonia synthesis
sys~em. The synthesis feed gas contains moisture and is
thereore Eed in~.o adsorber 202 in which essentially all
the water contained in ~he synthesis feed gas is removed
such that t~e total oxygen-containing compound content of
the synthesis feed gas is less than about lO ppmv. The
thusly dried gas is transported through ~he line 204 to
compressor 206 in which the synthesis feed gas is partially
co~pressed to su~stantially the superatmospheric synthesis
pressure, The partiall~ compressed synthesis feed gas is
compressed to a~ove the superatmosp~eric pressure in the
ammonia synthesis loop in compressor 208, Each of compressors
206 and 208 can be a multistage compressor, The effluent
from compressor 208 is passed via line 210 into the ammonia
synthesis loop where the synthesis feed gas is com~ined with
the gas circulating in the synthesis loop to provide the
reac~or feed gas. The reactor feed gas is compressed in
recycle compressor 2127 heated in heat exchanger 214, and
introduced into ammonla synthesis reactîon æone 216,
reaction effluent from the ammonia synthesis reaction zone
-
11~96~
-27- 07-0403A
is used as the ex~hange medium in heat exchanger 214 and is
cooled. The reaction effluent then passed to ammonia
coalescer 218 from which product ammonia is withdrawn via
line 220. Since the fresh synt~esis f~ed gas has not been
combined with the reaction effluent, a lesser weight of
gas needs to be refrigerated to condense the ammonia. The
overhead from ammonia coalescer 218 is recirculated in
conduit 222 (ammonia synthesis loop) to ammonia synthesis
reaction zone 216. Since the ammonia is removed rom the
synthesis loop prior to recompression in recycle compressor
212, less energy is expended in recirculating the gas in
the synt~esis loop.
A purge stream is withdrawn from the synthesis loop
via line 224, Since the purge stream is at a low temperature
because o~ the cooling to condense the ammonia product, the
purge stream is heated in heat exchanger 226 to suitable
temperatures for efecting the separation of hydrogen and
ammonia, e.g,, about 25 to 40C. The warmed purge stre~m
is passed to irst permeator 228.
First permeator 228 may be of any suitable design
including the design of the separator depicted in Figure 5.
The permeating gas exits permeator 228 via line 230 at
substantiall~ the pressu~e of the effluent from compressor
206. The permeating gas passes to and is com~ined with the
synthesis feed gas exitin~ compressor 206 in order to be
recycled to ammonia synthesis reaction zone 216. Since the
synthesis feed ga~ has been dried, the presence of ~mmonia
due to permeation through the separation membrane can be
tolerated in the synthesis feed gas to ~e compressed.
The non-permeating gas from first permeator 228 is
passed via line 232 to water scrubber 234 ~o remove ammonia.
Sincs a significant amount of the hydrogen has been separated
fr~m the purge stream in permeator 228, the water scrubber
can be of less volume than would ~e necessary if the
scru~ber were positioned upstream of the first permea~or,
The non-permeating gas having ammonia r~moved is passed
through line 236 to second permeator 238. A second
permeating gas is obtained ro~ second permeator 238 at a
112962~
-28- 07-0403A
pressure substantially the same as the pressure of the
synthesis feed gas in line 200. The second permeating gas
is passed through line 240 to line 200 whereat it is combined
with the synthesis feed gas in order to be recycled to ammonia
synthesis reaction 20ne 216. Since the second permeating gas
may contain water vapor from water scrubber 234, it is added
to the synthesis feed gas prior to the synthesis feed gas
being dried in adsorber 202~ The non-permeating gas from
second permeator 238 exits via line 242~
The ammonia synthesis system of Figure 4 is substantially
the same as the system depicted in Figure 3 except that the
purge stream in line 224 is compressed in compressor 22~ to
sufficiently elevated pressures that the first permeating
gas is at a pressure suitable for being directly introduced
back into the ammonia synthesis loop via line 231. Also,
the second permeating gas from second permeator 238 can be
at a higher total pressure than the second permeating gas
in the ammonia system depicted in Figure 3 even though the
total pressure dif~erentiaLs are essentially the same.
Accordingly, in the system of Figure 4, the second permeating
gas is combined with the synthesis feed gas downstream of
compressor 206. Instead of adsor~er 2Q2 as in the system of
Figur~ 3, coalescer 207 is provided between compressors 206
and 203 in order to r~move water vapor from t~e combined
synthesis feed gas and second permeating gas~
The following example is provided in illustration of a
process in accordance with th~^s invention~ ~11 parts and
percentages are b~ volume unless otherwise noted.
Ammonia is synthesi~ed from nitrogen and hydrogen
employing an ~mmonia synthesis plant similar to that depicted
in Figure 1. The hydrogen feedstock is obtained by primary
reforming of natural gas and the synthesis feed gas is
obtained by introducing air and the primary reformer effluent
into a secondary reformer. The effluent from the secondary
reformer is treated in a shift converter, a carbon dioxide
absorber and a methanator to provide approximately 52,000
kilograms per hour of a synthesis feed gas containing about
25 7 mole percent nitrogen, 73.1 mole percent hydrogen,
~.12962~
-29- 07-0403A
0.6 mole percent methane, 0.4 mole percent argon, and
O.2 mole percent wa~er. The synthesis feed gas is obtained
a~ abou~ 28 atmospheres absolute and 50C. The synthesis
feed gas is compressed to about 70 atmospheres absolute,
cooled to about 8C to condense water. The syn~hesis dried
feed gas îs ~urther compressed to about 133 a~mospheres
a~solute and îs introduced into and combined with the gas in
the ammonia synthesis loop. In the ammonia synthesis loop
the combined gases are compressed an additional 6 or 7
atmospheres and are treated in an ammonia coalescer which
removes a~out 44,500 kilograms o ammonia per hour. The
gases are heated to about 135 to 140C. Approximately
310,000 kilograms per hour of gas comprising about 66.5 mole
percent hydrogen, 22 mole percent nitrogen, 6.8 mole percent
methane, 3.5 mole percent argon, and 1,2 mole percent ammonia
are introduced into a Kellogg-type ammonia synthesis converter
utilizing a promoted iron $mmonia synthesis catalyst. A
reaction effluent gas at a temperature at about 280C is
obtained from the sgnthesis con~erter and contains about 11.4
percent ammonia. The effluent is cooled to a~out 43C. A
purge stream of about 2~1 percent of the gases in the synthesis
loop is removed, and the remaining gases are returned to the
s~nthesis loop compressor.
T~e purge stream is chil-led to a~out ~23C and about
1000 kilograms per hour of liquid ammonia are condensed and
removed ~rom t~e purge stream~ T~e purge stream contains
about 1,2 volume percent ammonia. The purge stream is then
scrubbed with water at about 25C at a water rate of about
2000 kilograms per hour. The purge stream contains less than
about 1~0 ppm~ ammonîa.
The purge stream is heated to about 3~G and then
passed to the first permeator which consists of 25 hollow
fi~er membrane-containing permeators in parallel. The
permeators are similar to tha~ depicted in Figure 5 and each
permeator contains about 93 square meters of ef~ective
sur~ace area. The m~m~ranes are comprised of anisotropic
polysulfone su~stantîally prepared in accordance with the
method disclosed in Example 64 of West German published
1~.2~62~
-30- 07-0403A
patent application DT 27 50 874 except that the spinning
solution contains about 30 weight percent solids; the
spinning jet dimensions are about 458 microns outside
diameter, 127 microns inside diameter, and 76 microns
diameter injection bore; the inj~ction fluid is a mixture
of 60 volume percent dimethylacetamide in water. The last
godet bath is at a temperature of about 50C; and the
fibers are washed for 24 hours with no subsequent storage
in water. Appropriate polymer solution and injection fluid
rates are employed such that the dimensions of the hollow
fibers are about 450 microns outside diameter and about 120
microns inside diameter. The permeator exhibits a separation
factor of hydrogen over methane of about 30 and a permeability
of about 50 x 10-6 cubic centimeters of hy~rogen per square
centimeters of surface area per second per centimeter of
mercury pressure drop. A pressure drop of about 65
atmospheres is maintained across the mem~rane, and
approximately 1100 kilograms per hour o a first permeating
gas is obtained from the bore side of the first permeator
stage. The first permeating gas comprises 90.3 volume
percent hydrogen, 6,2 volume percent nitrogen, 2,4 volume
percent methane, and 1.2 volume percent argon, The first
permeating gas is introduced into the feed gas exiting the
first compressor prior to the condensation of water from the
combined synthesis feed gas and first permeating gas stre~.
The non-permeating gas from the first permeator stage
is at a pressure of a~out 136 atmospheres absolute and
contains about 43.8 volume percent hydrogen, 35,4 volume
percent nitrogen, 13.7 volume percent me~hane, and 7.1
volume percent argon~ This non-permeating gas enters the
second permeator stage comprising 7 permeators Cas described
above2 in parallel. About 496 kilograms per hour of a
second permeating gas is obtained from t~e bore side of
the second permeator stage which is at a pressure of about
30 atmospheres absolute a~d comprises 88 volume percent
hydrogen, 7.4 volume percen~ nitrogen, 1.5 volume percent
argon, and 2.4 volume percent methane. The non~permeating
1129~26
-31~ 07~Q403A
gas from the second permeator stage is at a pressure of
about 136 atmospheres absolute and contains about 24.2
volume percent hydrogen, 47.7 volume percent nitrogen,
18.5 volume percent methane, and 9.6 volume percent argon~
About 86.3 percent of the hydrogen in the purge stre~m is
recycled to the 2mmonia synthesis reaction zone.
