ammonia traniee manual (general theory)

7/21/2019 ammonia Traniee Manual (General Theory)

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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.

ammonia traniee manual (general theory)

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