ammonia traniee manual (general theory)
TRANSCRIPT
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1.0 INTRODUCTION
Ammonia plant is based on steam reforming of natural gas. Ammonia plant consist of the
following sections
Natural gas supply
Desulphurisation section
Reforming section
CO conversion section
CO2 removal section
Methanation
rocess condensate stripping section
Ammonia synthesis section
Refrigeration section
Ammonia absorption section
urge gas recovery unit
!or simplified flow diagram refer schematic flow diagram
1.1 NATURAL GAS SUPPLY
Natural gas" which is used as feed stoc# and fuel for Ammonia lant and ower and
$team %eneration lant is supplied at Offsite %as Metering $tation by %as Authority of
&ndia 'td. from gas wells located in (rishna)%odavari *asin through pipe line. Natural
%as is supplied at a pressure of ++ (g,Cm2g. After metering in Offsites" natural gas for
process feed is directly received at Ammonia lant battery limit at +- (g,Cm2g and +-oC. !uel gas through separate header after pressure reduction is received at the battery
limit at (g,Cm2g. !uel gas is used for burners of !eedstoc# re)heater !)2-" rimary
Reformer !)2-/" Au0iliary $uper heater !)2-2 1 $tartup eater !)3-/ in Ammonia lant.
!eed gas received at battery limit goes to the desulphurisation unit for sulphur removal"
if any and subse4uently processed to produce synthesis gas for Ammonia production.
1.2 DESULPHURISATION UNIT
5he natural gas feed stoc# supplied to N!C' contains no 2$" but it is anticipated that
future supplies may contain sulphur compounds which have to be removed in order not
to poison the reforming catalysts and the '5 shift catalyst. Natural %as from battery limit
is heated to 63 oC in the !eed stoc# pre)heater !)2-" and is passed through the
ydrogenator. A bed of Nic#el)Molybdenum catalyst is provided to catalyse the
hydrogenation of organic sulphur compounds to hydrogen sulphide. 5here are two types
of organic sulphur compounds that may be present in the feed stoc#. One is called
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Reactive $ulphur7" containing 5hiophenes" 5hioethers etc. &n case of normal sulphur
e0cept Mercaptans ydrogen recycle gas is not consumed where as for less reactive
sulphur" recycle hydrogen is consumed as per the following hydrogenation reactions8
R$ 9 2 R 9 2$
:Mercaptans;
R/$R 9 22 R 9 R/ 9 2$
:5hioethers;
R/$$R 9 2 R 9 R/ 9 22$
:5hiophenes;
&f sulphur is present" natural gas is mi0ed with recycle gas from synthesis gas
compressor first stage discharge with flow of recycle gas around /-< NM,hr" in order
to avoid Carbon deposition on the catalyst due to catalytic crac#ing of higher
hydrocarbons if any. After preheating to 63 oC" the gas mi0ture passes to ydrogenator
Reactor R)2-/ and reacts to produce 2$. 5he above reactions are e0othermic but
insignificant :which depends on the type of $ulphur that determines the number of moles
of hydrogen ta#en up;. 2$ produced in R)2-/ and that already present in Natural %as is
then removed in 2$ Absorbers R)2-2 A,*" thereby the gas will be free of 2$. =ach
absorber contains one bed of >nO catalyst to absorb the sulphur. 5he absorbers are
operating in series with the second vessel acting as guard. ?hen the >nO in the first
vessel is getting e0hausted" a brea# through of 2$ from the first vessel may be
observed. 5he operation will then continue with the second vessel in service" while the
first vessel is being reloaded with fresh catalyst. 5he sulphur content at the e0it of R)
2-2* shall be less than -./ ppm on dry volume basis at all times which is tolerant to
reforming catalyst.
5he sulphur removal reaction in >nO bed ta#es place as follows8
>nO 9 2$ >n$ 9 2O
>nO 9 CO$ >n$ 9 CO2
>nO reaction with @$ depends on8
5ype of sulphur compounds.
5emperature8 &ncrease in temperature will generally increase the ability of >nO to
remove sulphur.
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Capacity8 As >nO reacts with sulphur it gets saturated with sulphur and looses its
activity. Normal life of >nO catalyst depends on the 2$ and sulphur concentration
in the natural gas.
1.3 REFORMING SECTION (GENERAL)
$team reforming is a vital part of the front end in plants producing Ammonia.
Developments in metallurgy have allowed steam reformers to be operated at higher and
higher levels of temperatures" pressures and heat flu0. 5he reforming process and the
design of the reformer are based on the reaction between methane and higher
hydrocarbons present in the natural gas with steam thereby generating CO" CO2 and
ydrogen. 5he ydrogen produced by the Methane reforming reaction is used to produce
Ammonia by combining with Nitrogen in the ratio of 8/ which is the stoichiometric ratio
of hydrogen and nitrogen to produce ammonia.
5he most important reactions ta#ing place in the reformer are
C+ 9 2O CO 9 2
CO 9 2O CO2 9 2
5hese reactions are ta#ing place in the presence of a nic#el)based catalyst. 5heoperating parameters maintained close to the e4uilibrium" are B
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!ollowing preheat" the gases are distributed through hair pin tubes into vertical reformer
tubes filled with Nic#el catalyst. 5he tubes are placed inside a !urnace" where sensible
heat and endothermic heat of reaction are absorbed in the tubes by radiation from a
number of wall burners to the tubes. 5he primary reforming of natural gas is done in a
5opsoe design side fired furnace" in comparison to the top fired furnace" where the
ma0imum heat input is concentrated in the top part of the furnace. &n the top fired
furnace during startup conditions with low flow" little or no heat of reaction in the tubes"
the ma0imum temperatures may well be found at the level of flames. &n such furnaces
higher than desirable temperatures may be present in the top part of the tubes even
when the outlet temperature is not higher than the level recommended.
&n the upper part of the top fired reformer" where the methane concentration is high and
hydrogen concentration is low" the potential for carbon formation is present. Due to the
radial temperature and concentration gradient in the tube" the ris# Eone e0tends
somewhat down along the hot tube wall. &f this Eone reaches a temperature level where
the rate of the crac#ing reaction becomes sufficiently high" carbon formation will ta#e
place resulting in a Fhot bandF. 5op fired furnaces are more prone to this #ind of
problem. Gsing the side)fired furnaces eliminates the above disadvantages of using top
fired furnace. &n case of side fired furnace" the reformer outlet temperature increases
gradually from the top towards bottom. 5he tube s#in temperature along the length of
the tube can be better)controlled in side fired furnace. 5he potential for carbon formation
with the age of the catalyst" the possibility of higher tube s#in temperature at the bottom
than from the top is better controlled using side fired furnace.
5he primary reformer furnace consists of /- tubes" inserted in two parallel chamberscalled the radiant Eone. =ach chamber has got 3 tubes in a single row. =ach row has
been divided into 3 sections. =ach section has got / tubes. Reformer tube outlets from
both chambers are connected to hot collectors through pigtails" which are placed outside
the rimary Reformer radiant Eone. ot collectors are again connected with cold
collector. 5he furnace operates with side firing of fuel gas on both sides of each row of
tubes to develop a process gas temperature of about B
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5he following reactions ta#e place simultaneously" producing a mi0ture of 2" CO" CO2"
C+and e0cess 2- when hydrocarbons undergo steam reforming over Nic#el catalyst8
Cnm 9 2n 2O n CO2 9 :2n9m,2; 2H eat :/;
C+ 9 2 2O CO2 9 + 2H eat :2;
C+ 9 2O CO 9 2H eat :;
CO2 9 2 CO 9 2- H eat :+;
Reactions start at 3-- oC for the higher hydrocarbons and
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available at higher temperature thereby facilitating the recovery of this heat by *!?
preheating.
ence" reforming pressure is fi0ed by an optimal balance between the reaction
e4uilibrium on one hand and compression power and heat recovery on the other. =4ually
important is the pressure drop across the reformer tubes. An increase in pressure drop
indicates possible catalyst fouling or partial bloc#age of tubes due to some other reason.
CARBON FORMATION
&n the operation of the primary reformer carbon may be formed partly outside the
catalyst" partly inside the catalyst. Carbon deposits outside the particle will increase the
pressure drop over the catalyst bed and deposits inside the particles will reduce their
activity and their mechanical strength. 5hermodynamically carbon formation is not
possible under the conditions foreseen" if e4uilibrium is obtained for each step. &f the
catalyst" however" is poisoned" e.g. by sulphur" it will loose its activity and carbon
formation is li#ely to occur. At very low steam to carbon ratio" there will be a possibility
of carbon formation" which would result in carbon deposits" especially inside the catalyst
particles. &f the catalyst is insufficiently reduced" or if it is partly o0idised during
production upsets" without subse4uent reduction" carbon formation may ta#e place.
Carbon deposition will hinder reforming and reduce heat transfer so that the tube wall
temperature will rise in that Eone producing 7hot bands7 and subse4uently 7hot tubes7.
recautions should be ta#en to prevent carbon formation on reforming catalyst for
successful reformer operations.
FLUE GAS SYSTEM
5he reformer furnace is designed to obtain ma0imum thermal efficiency by recovering
heat from the flue gases leaving the reformer radiant section. 5he hot flue gas from top
at 6- oC passes through downward and horiEontal flue gas duct. 5he flue gas blower :()
2-/; induces the desired draft of B3 MM ?C. 5he flue gases enter the waste heat
recovery section and give up heat successively to the various coils. At the outlet" the flue
gas temperature is reduced to appro0. /B- oC as any further reduction in temperature
may result in condensation of sulphur compounds if any present in the flue gas.
COMBUSTION AIR
Combustion air to the forced draught radiant burners of primary reformer and au0iliary
steam super heater is supplied by Combustion Air *lower" ()2-2 after preheating to 2oC in combustion air pre)heater =)2-+ by recovering sensible heat of the flue gases.
SECONDARY REFORMING
5he partially reformed gas e0it rimary Reformer contains /+.- mole of C + :dry
basis;. 5he methane content is further reduced to -.< mole percent :dry basis; at hightemperature in the secondary reforming step. &n the $econdary Reformer" R)2-" the
heat is supplied by combustion of part of the gas achieved by mi0ing air into the gas as
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compared to the indirect heat by firing in the rimary Reformer. 5his combustion
provides heat for the rest of the reforming in R)2-. 5he methane slip e0it rimary
reformer is so adIusted that the process air supplying the reaction heat in the $econdary
Reformer will give the ydrogen, nitrogen ratio of 8/ in the syn. gas. &t is desirable to
reduce the methane content of the process gas to a low level in order to #eep the level
of inert gases low. 5he methane content e0it R)2- is dependent upon the methane slip
at the rimary Reformer outlet at specified conditions. 5he high C +slip at !)2-/ outlet
gives rise to C+slip of -.< mole :dry; at R)2- outlet at +oC. $ince air 4uantity is
fi0ed when %R Gnit is running" the 28N2ratio in the ma#e up gas is 2.B6 which gives
rise to 8/ at Ammonia $ynthesis Converter inlet by recovering 2from %R. 5he inert
concentration is maintained at 6 in $ynthesis loop at Converter inlet.
5he partially reformed gas from primary reformer is directed to the refractory lined
$econdary Reformer R)2- at / (g,Cm2g and B
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Reformed gas with unreacted process steam at + oC from the bottom of $econdary
Reformer R)2- passes through the tube side of special type of ?aste eat *oiler =)2-6.
5his ?aste eat *oiler consists of two compartments" which is castable refractory lined.
5he total gas is passed through tube side of the first compartment where as the second
compartment has been provided with internal bypass to control e0it temperature. A
temperature controller regulates flow through the bypass to maintain 5 $hift Converter
inlet temperature at about
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through the catalyst bed. 5he outlet temperature is +2+ oC. At the main startup of the
plant" the catalyst must be activated" which is mainly a reduction of !e2- to !e-+. 5he
reduction will ta#e place at a temperature above 23- oC" but the temperature should not
be allowed to e0ceed +-- oC during the reduction in order not to decrease the activity of
the catalyst. ?hen new" the catalyst can operate at a gas inlet temperature of 3- oC.
Afterwards" the optimum inlet temperature will be higher" but as long as the outlet
temperature has not reached +B- oC" the activity will only decrease slowly.
5he cold catalyst can be heated by steam alone" both when o0idiEed and reduced. Drops
of li4uid water on the hot catalyst may result in disintegration of the catalyst. 5he
catalyst is very sensitive to salts" which may be introduced with the steam. 5he content
of chlorine in the gas should be well below -./ ppm. 5he catalyst is not influenced by
sulphur in the 4uantities present in this plant. 5he fresh catalyst contains" however"
about -. sulphate" which will be given off as 2$ during the first wee# of operation.
Normally the catalyst is not o0idiEed by steam alone" but should be o0idiEed by adding a
small amount of air to the steam before it is accessible as it is pyrophoric in reduced
state.
HEAT RECOVERY FROM HT SHIFTED GAS
5he final shift reaction is completed in 'ow temperature shift converter R)2-3. 5he gas
leaving 5$ is cooled to 2--
o
C before entering '5 shift converter by recovering wasteheat successively in ?aste eat *oiler after Co)converter =)2/- and *!? re)heater =)
2//A,* and 5rim eater =)2-. &n =)2/-" ($ steam is generated while cooling the gas to
+- oC. art of gas is sent to the 5rim eater =)2- to preheat the Methanator feed inlet
gas partly. 5he gas then passes through the shell side of =)2//A,* and gets cooled to
2-- oC. 5here is a bypass of =)2//A,* for controlling '5 $hift Converter gas at desired
inlet temperature.
LO TEMPERATURE SHIFT CONVERTER
5he '5 $hift Converter R)2-3 contains 6-. M of the catalyst consisting of o0ides of
copper" Einc and aluminium. As the catalyst is e0tremely sensitive to sulphur which may
be liberated not only from the preceding 5 shift catalyst but also from secondary
reformer refractory material" the '5 shift converter is bypassed during initial stage until
the gas is practically sulphur free. 5he chlorine may be present in process steam and
4uench water" due to mal)operation of the water treatment system and process air due
to atmospheric air pollution in very small amounts. *esides chlorides and sulphur"
gaseous $i compounds are also catalyst poisons. ?hen the catalyst is in a reduced state"
temperatures above 23- oC must normally be avoided. A short e0posure to -- oC will
have no adverse effect on the catalyst. Normal operation should ta#e place at as low atemperature as possible. ow ever" at temperatures near the dew point" the activity will
decrease because of capillary condensation of water inside the catalyst thus reducing
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the free area. During operation" the temperature should" therefore" be #ept at least 2-oC above the dew point of the gas. 5he reduced catalyst is pyrophoric and has to be
o0idiEed before opening of the converter. 5he normal operating temperature is between
2-- oC and 2/6 oC.
5he actual temperature of the inlet gas to R)2-3 to be selected is dependent on the
activity of the catalyst.
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1.! CO2 REMOVAL SECTION
5his unit provides process gas free of CO2 :limit /--- ppm; for the production of
ammonia and necessary CO2 for Grea production. &n this unit" CO2 in the process gas is
absorbed by the %J solution in the Absorber" C)-/ thus providing process gas with less
than /--- ppm of CO2. $tripping of the absorbed CO2 is done in the two regenerators
and CO2 stripped is supplied to Grea lant. CO2 removal section #now how is by
%iammarco)Jetroco#e of &taly. 5he Jetroco#e solution consists of (2 CO" Janadium
ento0ide" glycine and D=A where J2O3 :Janadium ento0ide; is the corrosion inhibitor
and glycine,D=A are the activators. 5he chemistry involved in this unit is chemisorption
and is e0plained as follows8
CO2 9 2O CO) 9 9 :/;
(2CO 9 CO) 9 9 2(CO :2;
(2CO 9 CO2 9 2O 2(CO :;
5he reaction rate of :; depends on the reaction rates of :/; and :2;. Reaction rate of :/;
is slow and the activator activates this reaction by 4uic#ly introducing the gaseous CO2
in the li4uid phase. 5he activator glycine reacts with CO2 and forms glycine carbonate
according to the reaction.
N2C2COO) 9 CO2 COO
)N C2COO) 9 9 :+;
COO)N C2COO) 9 2O N2C2COO
) 9 CO) :3;
5he sum of :+; and :3; gives :/;.
&n solution regeneration" reaction :; is reversed by application of heat and pressure
reduction and the lean and semi)lean (2CO solution is recirculated for further
absorption of CO2. 5he process gas from J)2-6 enters the CO2 removal section at 2B.3
(g,cm2g and /
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water preheaters =)-
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separator J)-2. 5he CO2is fed to the *ooster compressor ()-/ or it can be vented to
atmosphere through &C)-2
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5he main reason why the reaction is reversed is the lower temperature favoring
formation of methane. Other critical variables governing the reactions are pressure and
steam content. owever" within the allowable temperature range" the e4uilibrium
conditions are so favorable that practically only the catalyst activity determines the
efficiency of the methanation. 5he higher the temperature" the better the efficiency" but
at the same time it means a shorter lifetime for the catalyst.
!urther more in case of a possible brea# through of CO2 and CO to the methanator
which would result in a higher temperature rise" a low inlet temperature is preferred as
this limits the temperature rise. After the methanator the gas normally contains /- ppm
of CO 9 CO2. 5he temperature rise of gas in methanator will normally be about 2/ oC.
Methanator contains 2< M of catalyst and has appro0imately the same characteristics
as of reformer" being nic#el catalyst on a ceramic base. As the reactions ta#e place at
much lower temperature than those prevailing in the reformers" the catalyst must be
very active at low temperatures. 5he catalyst is sensitive to Arsenic" $ulphur and
Chloride Compounds. 5he adiabatic temperature rise per mole of CO is B+ oC and per
mole of CO2 is
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NIC#EL CARBONYL GAS
Nic#el carbonyl gas is a poisonous and to0ic gas" which may be present in R)-/. Gnder
certain conditions" CO in the process gas reacts with the catalyst Ni to form Nic#el
carbonyl gas.
+ CO 9 Ni Ni:CO;+
5he favorable temp. range of the formation of this gas is between +3 oC and 2-3oC.
ence R)-/ catalyst should never be allowed to cool in the presence of CO containing
gas. Rather it should be purged out with N2at the time of shut down. ?hile heating the
catalyst with process gas containing CO" heating should be done faster in the range of +3
oC to 2-3 oC.
1.$ PROCESS CONDENSATE STRIPPING SECTION
rocess condensate from J)2-6" J)-/ and J)// contains small amounts of dissolved
CO2" Nand Methanol. CO2 removal section O condensate from J)-2 and J)-
contain mainly dissolved CO2. 5he dissolved gases in the condensate from the above
separators are to be reduced to a tolerable limit before sending it to DM lant for further
treatment. 5he condensate from J)-2 and J)- at +- oC is pumped by ) -3A,* to a
flash separator J)32" which is a pac#ed column and operating at atmospheric pressure.
5he condensate enters the vessel on the top of the pac#ed bed. Condensate from J) //
at +/ oC" from J)-/ at //.3 oC and from J)2-6 at /
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processing. &n this way rocess Condensate pollution in li4uid effluent is totally
eliminated.
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1.% AMMONIA SYNTHESIS SECTION
5he Ammonia $ynthesis ta#es place in the Ammonia converter R) 3-/ as per the
following reaction.
&ron Catalyst
N2 9 2 2N 9 eat
5he reaction is limited by the e4uilibrium concentration and only part of the
ydrogen and Nitrogen can be converted into Ammonia per pass through the
Catalyst bed . 5he e4uilibrium concentration of Ammonia is favored by high pressure
and low temperature. owever" reaction rate is very much enhanced by high
temperature operations. 5here is a compromise between thermodynamic e4uilibrium
1 reaction #inetics. As a result there is an optimum level for the Catalyst
temperatures at which the ma0imum production is obtained. At higher temperatures
the e4uilibrium percentage :which is the theoretically highest obtainable
concentration of Ammonia; will be too low while at lower temperature the reaction
rate will be too low. 5he $ynthesis loop is designed for a ma0imum pressure of /33
(g,Cm2g and the normal operating pressure is in the range of //)/+/ (g,Cm2g.
5he reaction temperature in the catalyst bed is
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o0ide to free &ron. 5his reduction is carried out with circulating $ynthesis gas. 5he
Catalyst activity will decrease slowly during normal operation and the lifetime of
Catalyst is 6 to /- years. 5his is again influenced by the actual process conditions.
Notably the temperature in the Catalyst bed and the concentrations of Catalyst
poisons in the $ynthesis gas at converter inlet. $ulphur compounds and compounds
containing O0ygen such as water :2O;" Carbon Mono0ide :CO; and Carbon dio0ide
:CO2; are all poisons to the Catalyst and small amounts of the catalyst poisons will
cause a considerable decrease in Catalyst activity. art of the poisoning effect is only
temporary and catalyst activity will recover somewhat when the gas is clean again. A
certain permanent decrease in the Catalyst activity will however remain and high
concentrations of O0ygen compounds at converter inlet even for short duration
should therefore be avoided.
PROCESS CONDITIONS
Ammonia $ynthesis reaction is affected by the following parameters8
Ammonia content in the feed gas
&nert gas content in the feed gas
2to N2ratio in the feed gas
Reaction temperature
Circulation Rate
Operating pressure
Catalyst activity
AMMONIA CONTENT IN THE FEED GAS
A low Ammonia concentration at converter inlet gives a high reaction rate and thus a
high production capacity. 5he Ammonia concentration at converter inlet is dependent
on the cooling level in the refrigeration chillers and the operating pressure. At
converter inlet +./ Ncorresponds to )3oC at a pressure of /2 (g,Cm2g in the
Ammonia $eparator" J)3-/.
INERT GASES
5he Ma#eup gas contains /. :Jol.; of argon and methane. 5hese gases are inertsin the sense that they pass through the $ynthesis converter without undergoing any
Chemical changes. *ut a high concentration of inerts reduces the partial pressures of
ydrogen and Nitrogen thereby reducing the conversion. 5herefore a constant purge
of gas from the loop is maintained to #eep the inerts level in the converter inlet at
about 6. 5he catalyst activity decreases with the catalyst age. 5his can be
compensated by either increasing the loop pressure and the circulation rate or by
decreasing the inert level.
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HYDROGEN/NITROGEN RATIO
*y the $ynthesis reaction" volumes of ydrogen react with / volume of Nitrogen to
form 2 volumes of Ammonia. 5herefore the 2,N2ratio in the loop and ma#eup gas
must be close to 8/. A small change in 2,N2ratio of the ma#e up gas will result in a
much bigger change in the 2,N2ratio of the circulating $ynthesis gas. 5he 2,N2
ratio of the ma#eup gas should normally be about 2.B6 so that after addition of
recovered hydrogen from %R Gnit the ratio will be about .-. 5he $ynthesis loop is
designed for operating at the 2,N2ratio of .-" but special conditions may ma#e it
favorable to operate the loop at a slightly different ratio in the range of 2.3 to .3.
?hen the ratio is decreased to 2.3" the reaction rate will increase slightly :but fall
again for ratios below 2.3" while on the other hand" the circulating $ynthesis gas will
be heavier. 5herefore the pressure drop through the loop will increase and the
Ammonia separator efficiency may decrease" leading to increased Ammoniaconcentration at the converter inlet. 5he 2,N2ratio in the loop should be #ept as
constant as possible. 5he ratio is controlled by the 2,N2ratio in the ma#eup gas
which will have to be adIusted to get desired ratio in the circulating gas. After
ma#ing any change in the 2,N2ratio of the ma#eup gas sufficient time should be
allowed for the system to find its new e4uilibrium before ma#ing further changes.
REACTION TEMPERATURE
5he temperatures in the Catalyst bed are usually in the order of
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appro0. 6- oC before entering the 2nd bed. &n the 2nd bed the gas outlet
temperature is about + oC.
CIRCULATION RATE
5he capacity of the synthesis loop with regard to Ammonia production rises with
increasing circulation rate. owever" the Ammonia production per cubic meter of
circulation gas which is proportional to the temperature difference between converter
e0it and converter inlet" will decrease.
OPERATING PRESSURE
5he $ynthesis loop is designed for a ma0imum pressure of /33 (g,Cm 2g and it is
foreseen that the $ynthesis loop can operate at a pressure of /+2 (g,Cm 2g when
operating at design production rate" design inert level and design gas composition.
5he actual operating pressure is not directly controlled and is dependent on the otherprocess conditions" notably production rate" inert level" ammonia concentration at
converter inlet" 2,N2ratio and Catalyst Activity. 5he production rate increases with
rising pressure and for a given set of process conditions" the pressure will adIust
itself so that the production rate corresponds to the amount of Ma#eup gas fed into
the loop. 5he loop pressure will be increased by increasing the Ma#eup gas flow to
the loop" by decreasing the circulation rate" increasing the inert level or the
concentration of Ammonia at converter inlet and by changing the 2,N2ratio away
from the optimum. 5he decreases in Catalyst activity will also increase the operating
pressure.
1.& REFRIGERATION SECTION
5he refrigeration section is used to li4uefy gaseous Ammonia and consists of a
compressor unit" a condenser" an accumulator and number of chillers. 'i4uid N
flows from the Naccumulator" J)3-6" to the Nheater =)3/" which serves to
heat Ammonia sent to the Grea lant. 5hen the li4uid Nflows partly to ma#e up
gas chiller" =)3- partly to the /st Nchiller" =)3-3 and partly to the inert gas
chiller" =)3//. &n =)3-3 1 =)3- the li4uid N is evaporated at
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1.10 AMMONIA ABSORPTION
Ammonia produced in the Ammonia $ynthesis Converter" R)3-/ is separated from
the unreacted gas mi0ture in Ammonia $eparator" J)3-/. Due to high pressure and
low temperature in J)3-/" the loop gases are dissolved in li4uid ammonia to a
certain e0tent depending on their respective solubility. 5hese gases are removed
from li4uid ammonia by successfully flashing the Ammonia li4uid from /2 (g,Cm 2g
to 23 (g,Cm2 g in the 'etdown vessel J)3-2 and further to near atmospheric
pressure in the !lash vessel" J)3-. 5he vapours from the 'etdown vessel" J)3-2
which contain /3.3 :Jol.; Ammonia are sent to Ammonia Absorption unit to
recover Ammonia. ?hile the vapours from flash vessel" J)3- are sent to
Refrigeration Compressor" ()+3/ and after compression along with other Ammoniavapours from the refrigeration circuit are condensed in N condenser and sent to
Ammonia accumulator" J)3-6. 5he uncondensed gases from Ammonia Accumulator
J)3-6 are again cooled in &nert gas chiller =)3// for recovering some Ammonia and
the inert gases which also contain about +6.3/ :Jol.; Nare sent to Ammonia
absorption unit to recover balance Ammonia. 5he inert and 'etdown gases are sent
to the bottom of Ammonia absorber" C)322 where they are contacted with water
flowing down in the pac#ed beds. 5he Ammonia is absorbed in water while the gas
mi0ture containing 2" N2" Ar" and C+is sent as off gas to be burnt in Au0iliary
steam super heater" !)2-2. 5he Ammonia water mi0ture along with a4ueous N
from %R is sent to the distillation column" C)32/ where Ammonia is separated fromwater. 5he water is again circulated to the absorber" C)322 and partly to the absorber
in %R unit. 5he overhead Ammonia gases from the distillation column are
condensed in Condenser =)322 and Nli4uid is sent to J)3-.
1.11 PURGE GAS RECOVERY UNIT
&n order to maintain the inerts :C+9 Ar; concentration in the $ynthesis loop at a
constant value a continuous purge is ta#en from the downstream of 2nd cold eat
=0changer" =)3-6 where the inerts concentration is the ma0imum. 5ypically purge
gas is about
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operation due to C+slip from secondary reformer and CO slip from '5$ converter is
greatly reduced. &n the %R unit cryogenic cooling in cold bo0 separates hydrogen"
where C+" Ar and most of N2are condensed. $o more process Air :N29 O2; must be
introduced to secondary reformer in order to maintain 28N2 ratio as 8/ in the
$ynthesis loop. 5he increased rocess Air results in greater combustion in $econdary
Reformer" which means" more C+ slip from rimary Reformer can be allowed.
Conse4uently" rimary Reformer duty is decreased marginally" resulting in better
tube life and overall less energy consumption.
5he purge gas is sent at / (g,Cm2g and / oC to %R Ammonia absorber C)33-"
where Ammonia is absorbed in water. 5he Ammonia ) water mi0ture :A4ueous
Ammonia; from absorber C)33- is sent to the Distillation column" C)32/ for
separation of Ammonia. 5he urge gas" which contains traces of Ammonia and water
is letdown to B- (g,Cm2g and sent through the adsorption unit where molecular
sieve is used to adsorb traces of ammonia and water. 5he dry gas now containing
only 2" C+" N2and Ar is sent to the cold bo0 where the gas is cooled to below )/6-oC where almost all the C+ and argon and some of the N2 are condensed. 5he
uncondensed ydrogen along with some 4uantity of Nitrogen and small percentages
of C+ and Argon is sent bac# to the rd stage suction of the $ynthesis gas
compressor" ()+/.
5he gases while leaving the cold bo0 cool the incoming feed gas. 5he low
temperature in the cold bo0 is maintained by the flashing of condensed C+to 3
(g,Cm2g which is removed as off gas and along with off gas from Ammoniaabsorber" C) 322 is used for burning in Au0iliary $team $uper heater !)2-2.