BRIEF DESCRIPTION OF AMMONIA & UREA PLANT WITH REVAMP

Author

Prem Baboo

Sr. Manager (Prod)

National Fertilizers Ltd. India

An Expert for www.ureaknowhow.com

Fellow of Institution of Engineers India

Specialist for Heat Exchangers Design

1.1 Company Profile

National Fertilizers Limited, a Govt. of India Undertaking, was incorporated on

23rd August 1974. It is the second largest producer of nitrogenous fertilizer in

the country and has four operating fertilizer units located at Nangal, Bhatinda,

Panipat and Vijaipur with a total installed capacity of 32.083 lakh tones Urea.

All the fertilizer units of NFL are operating and in good condition.

The Vijaipur unit, which is an ISO 9001:2000 & 14001 certified, comprises of

two streams-Vijaipur-I and Vijaipur-II, which went into commercial production

in July, 1988 and March, 1997 respectively. The Complex comprises two

streams of Ammonia Plants each of 1750 MTPD & 1864 MTPD for line-I & line-

II respectively capacity (after revamped) based on “Steam reforming process” of

Haldor Topsoe, Denmark and four steams of Urea Plants each of 3030 &

3231TPD MTPD capacity for line-I & line-II (after revamped) based on

“Ammonia stripping process” of SAIPEM (erstwhile Snamprogetti),Italy and

together with all necessary utilities and infrastructure facilities. Vijaipur II is

more energy efficient than Vijaipur-I Plant due to incorporation of energy

saving Equipments & processes at the design stage itself. It is also having

provision of mixed feed (Natural Gas + Naphtha) thus giving more flexibility of

operation. Both the plants have consistently achieved high levels of capacity

utilization.

Site location

The plant site is about 30 km from the District town Guna along the National

Highway (NH-3) towards Indore. The nearest railway station is Ruthiai junction

at a distance of 6 km. A broad-gauge line connects Bina (150 km) from the East

and Kota (170 km) from the Northeast to Ruthiai. From here, another broad-

gauge line runs to Maksi 240 km away which in turn connects Bhopal and

Ratlam. Very close to the site, the rain-fed river “Chopan Nallah” flows and

another river “Parvati” is 6 km away from the site.

Infrastructure facilities

The infrastructure facilities comprises the gas pipeline, water supply, power

supply from state grid through transmission line laid by MP State Electricity

Board, Railway siding and communication facilities. These facilities are already

available at the existing plant site and only augmentation of some of the

facilities, if required, will be done.

Raw water

The existing Vijaipur fertilizer complex has permanent raw water supply

facilities from Gope-Krishna Sagar Dam nearly 15 km away from the site.

Power

At Vijaipur unit, three Gas Turbine Generator (GTG) sets each of 16 MW rating

(at ambient conditions) have been provided. The exhaust gases from GTG are

used in HRSG for generating steam. At present, two GTGs are running at around

13 MW load each.

Railway Siding

After implementation of the proposed debottlenecking scheme, the existing

railway infrastructure can handle the total traffic.

Social Infrastructure

The existing fertilizer complex has well-established residential colony

(Township) with all common facilities such as hospital, school, shopping centre,

recreation centre, guesthouse and the sports facilities etc.

Background of the Proposal

Chemical fertilizers have played a vital role in the success of India's Green

Revolution and consequent self-reliance in food-grain production. The increase

in fertilizer consumption has contributed significantly to sustainable

production of food grains in the country. The Government of India has been

consistently pursuing policies conducive to increased availability and

consumption of fertilizers in the country. The Government's policy has hence

aimed at achieving the maximum possible degree of self-sufficiency in the

production of nitrogenous fertilizers based on utilization of preferred feedstock

as NG/RLNG. Presently, various fertilizer companies of the country are

planning for augmentation of their capacity of existing ammonia-urea plants

based on the study report conducted by respective process licensors/ Indian

consultant. NFL Vijaipur is also contemplating in the same line.

1.5 Need for the Revamp

The volatile international urea fertilizer market could be reason for country's

strategic policy. This is also desirable as the international market is very

sensitive to demand supply scenario. The international price of urea is mainly

governed by the demand potential of two countries i.e India and China. The

international price of urea is governed not by the cost plus profits approach but

by ‘opportunistic’ considerations. The international price of urea at present is

high and the scenario is likely to continue.

1.6 With the above consideration, NFL has initiated for capacity augmentation as

well as energy saving measures of existing Vijaipur-I & II ammonia & urea

plants. The rationale behind the proposed revamping project could be:

a] Availability of existing infrastructure facilities.

b] Lower capital cost, cost of production and retention price compared to new project consequently lower incidence of subsidy.

c] To meet the gap in demand and supply in the country.

1.7 The specific energy norms fixed by FICC was 6.271 Gcal and reduced to 5.952

during stage-II pricing. In present scenario, NG supply is lean in nature which

makes the plant unable to convert total ammonia produced to urea due to

shortage of CO2. In case of lean gas, urea plants will not be in a position to

produce urea even at reassessed capacity. In view of above and for the survival,

NFL has got the feasibility study done by M/s. HTAS, SAIPEM and PDIL. The

study identified certain energy saving schemes. NFL is in the process for

implementation of some of these schemes. Due to implementation of these

energy saving measures, there will be capacity enhancement of

Ammonia-I of nearly150 MTPD & the plant will be able to produce 1750

MTPD(with PGR),

Ammonia-II nearly 225 MTPD & the plant will be able to produce 1864

MTPD (without PGR).

Urea-I plant will be able to produce 3030 MTPD

Urea-II plant will be able to produce 3231MTPD

To meet the shortfall of 357 MTPD CO2 in achieving the capacity of Urea

plants as stated above, a 450 MTPD Carbon-Dioxide Recovery (CDR) Plant

is to be installed.

PROCESS DESCRIPTION

2.1 Ammonia Plant

The ammonia process is based on Haldor Topsoe Technology.

The process steps involved in production of ammonia are: • Desulphurization and Reforming • Carbon Monoxide Conversion • MDEA Carbon Dioxide Removal • Methanation • Ammonia Synthesis Loop • Ammonia Refrigeration • Ammonia Recovery • Process Condensate Recovery • Steam System

The descriptions of the various process steps are as follows:

Desulphurization

Natural gas feedstock containing minor quantity of sulphur compounds are

required to be removed in order to avoid poisoning of the reformer catalyst in the

primary reformer and the low temperature shift catalyst in the CO converter. The

low temperature shift conversion catalyst, in particular, used in LP converter, is

sensitive to deactivation by sulphur and sulphur bearing compounds. Natural gas

from the battery limit, after compression at 45 kg/cm2 is mixed with recycle gas

containing H2 and heated to 400oC in the waste heat recovery section of primary

reformer. Hot feed natural gas along with recycle H2 enters hydrogenator, filled

with hydrogenation catalyst where organic sulphur compound are converted to

H2S. H2S is absorbed on a specially prepared zinc oxide catalyst, contained in

sulphur absorber.

The sulphur contained in the feedstock will be reduced to a very low level i.e. 0.05

to 0.1 ppm sulphur by weight.

Reforming

The reforming of the hydrocarbon feed takes place in two stages - first in a direct

fired primary reformer and later in an auto-thermal catalytic secondary reformer.

The hydrocarbon feed coming from the desulphurization unit is mixed with steam.

The steam/carbon ratio is 3.0. The reaction mixture is preheated in coil bank

located in waste heat recovery section of primary reformer furnace before entering

the catalyst tubes of primary reformer, where it is converted into hydrogen, carbon

monoxide and carbon dioxide by reaction inside the catalyst filled reformer tubes.

In the secondary reformer, preheated process air is added and the heat thus

generated by burning of hydrogen of the reformed gas is used for supplying heat

required for conversion of residual methane coming from primary reformer. The

methane concentration in the outlet gas from the secondary reformer is around 0.3

vol% (dry basis).

The reforming unit consisting of a primary reformer with a waste heat section, and

a secondary reformer, is briefly described below:-

Primary Reformer

The primary reformer which is of side-fired type consists of two chambers. The

chambers are placed side by side and functions as one unit. The two furnace

chambers are joined to a common flue gas duct at the top which is connected to

flue gas heat recovery section, housing heat recovery coil banks located a t

different levels from top to bottom.

Each furnace chamber contains a number of vertically mounted micro alloy

reformer tubes filled with primary reforming catalyst. The tubes are mounted in

a single row along the centre line of the chamber. The process gas flows

downwards with the gas being distributed to the top of the tubes from a header

through "hairpins" at a temperature of about 520oC. The gas leaves the tubes

through bottom "hairpins" and enters a refractory lined collector connected to

number of symmetrically placed high alloy hot collectors.

The tubes are heated by a number of burners located in each side wall of the

furnace chambers and arranged in horizontal rows at several elevations to

provide easy control and maintain uniform temperature profile along the length

of the catalyst tubes. In this manner, the optimal utilization of the expensive high

alloy tubes is obtained. The flue gas flow is upwards with outlet near the top of

the radiant chamber. The flue gas outlet system comprises of a common flue gas

collector mounted between the two radiation chambers. The flue gas

temperature going out of radiant chamber is about 1100oC.

Flue Gas Heat Recovery Section

The flue gas passes via the flue gas duct to the flue gas heat recovery section, in

which the sensible heat of the flue gas is utilized for the following duties. • Preheating of hydrocarbon/steam mixture going to the primary reformer, • Final preheating of process air for the secondary reformer, • Final superheating of HP steam, • Final preheating of natural gas, • Superheating of HP steam, • Preheating of process air for the secondary reformer, • Preheating of natural gas, • Preheating of combustion air,

At the outlet of the flue gas duct, the flue gas temperature is reduced to

approximately 150oC-165oC. An induced draught fan is used for discharging the

flue gas to atmosphere by venting continuously through a chimney stack. Secondary Reformer

The gas from the primary reformer passes to the Secondary Reformer through a

refractory lined transfer line. The gas is admitted to the vessel through a top

dome-mixing chamber, where it is mixed with the process air, which has been

compressed to 38 kg/cm2g in the air compressor, and preheated to around 550oC

in the flue gas heat recovery section. The secondary reformer is a refractory

lined vessel. The burner mixer is mounted at the top of the vessel. The vessel

contains a bed of a nickel reforming catalyst, supported by a grid of high

temperature resistant material.

The balance between the reforming reaction taking place in the primary and

secondary reformers depends to a great extent on the preheat temperatures and

the methane leakage. In practice, the firing in the primary reformer is adjusted so

that the desired outlet conditions from the secondary reformer are obtained with

the amount of process air required to maintain a hydrogen/nitrogen ratio of

approximately 3 to 1 in the make-up synthesis gas going to ammonia synthesis

loop.

The high temperatures in the primary and especially in the secondary refor mer

necessitate chemical resistivity of the catalysts to the constituents of secondary

reformer lining material. Particular emphasis is given with regards to the use of

catalysts free from silica and alkali, together with the use of refractory lining

material having a very low content of silica and iron. Presence of silica and iron

in higher proportions induces formation of volatile compounds which are easily

carried out of secondary reformer and deposited on the surfaces of waste heat

boiler tubes.

The process gas leaves the reforming section at about 1000oC - 1020oC. It is

cooled to about 392oC-400oC in the waste heat boiler, where 120 kg/cm2g

saturated steam is produced. The process gas is further cooled to 360 oC in the

steam super heater. After cooling, the gas flows to the high temperature CO

converter.

CO Conversion

The CO conversion takes place in two adiabatic stages.

The high temperature CO converter contains a Cu-promoted high temperature

shift catalyst. High activity, high mechanical strength and very low sulphur are

the main characteristics of this variety of catalyst. The low temperature CO

converter is loaded with low temperature shift catalyst, which is characterized,

by high activity, high strength, and high tolerance towards sulphur poiso ning. A

top layer of a special catalyst generally termed as the guard catalyst is installed

to absorb any possible chlorine carry over in the gas and also to prevent liquid

droplets from reaching the main bed of LT shift conversion catalyst. After reforming, about 13-14% CO is present in the gas (dry basis). In the high temperature CO converter, the CO content is reduced to approximately 3.3 vol%, and in the process, the temperature of the product gas is increased from 360oC to 435 oC. HT shift reactor effluent gas is cooled in stages to around 200 oC -210 oC before entering low temperature CO converter, in which the CO content is reduced to approximately 0.3 vol %, while the temperature of the product gas increases to 228 oC.

The heat content of the effluent from the high temperature CO-converter is

recovered in the trim heater, in the high pressure waste heat boiler, and in high

pressure boiler feed water preheater.

CO2 Removal

The gas leaving the CO conversion unit contains a considerable amount of

recoverable heat. Owing to the steam content of the gas mixture, this heat is

present mainly as latent heat. The waste heat in process gas is recovered in a

high pressure BFW preheater, in the stripper re-boiler, and in the de-mineralized

water preheater, E 305.

For the removal of CO2, GV process for Ammonia-II and Benfield process for

Ammonia-I are present. Main equipment in the process is the CO2 absorber, and

the CO2 stripping columns. The absorbent solution consists of an activator. The

regenerated CO2 will be available at 40 oC and 0.15 kg/cm2g pressure, from LP

flash cooler KO drum. An electric motor driven CO2 booster compressor shall be

used for delivering CO2 at a pressure of around 0.5 kg/cm2g at the battery limit

of urea plant.

Bulk quantity of carbon dioxide is removed from the process gas in the lower

part of the absorber, by scrubbing with flash-regenerated solution coming from

LP flash column. In the upper part of the absorber, strip-regenerated solution

generated in CO2 stripping column, is used for scrubbing.

The flash regeneration of the rich solution is performed in two stages. In the HP

flash drum, a large part of the dissolved inert components is expelled at a

pressure of around 8 kg/cm2g. The flashed gas is then compressed and recycled

to the CO2 absorber, for maximum recovery of the available CO2. The flashed

solution from the bottom of the LP flash drum is divided and the major part of

the solution is transferred to the lower part of the absorber. Only a minor part of

the flashed solution is fed to the CO2 stripping column, where CO2 is stripped out

with steam to obtain a low residual CO2 loading. The lean solvent from the

bottom of the stripper is pumped to the top section of the CO2 absorber.

In this way, nearly complete removal of CO2 can be achieved with de-carbonated

gas containing only 0.05 vol% CO2 (on dry basis) at the expense of very low heat

consumption.

Methanation

After CO2 removal, the de-carbonated gas contains 0.05% CO2 and 0.4% CO (dry

basis). These compounds are poisonous to the ammonia catalyst and must be

removed before the gas is pushed to the ammonia synthesis loop. This is

accomplished in the methanator, where CO and CO2 react with H2 to form CH4,

which is harmless to the ammonia catalyst. The reaction takes place over a

nickel-based catalyst. The content of CO+CO2 is reduced to less than 5 ppm.

The inlet temperature to the reactor is maintained at around 300oC, and the

outlet temperature around 325 oC. The inlet gas is preheated by heat exchange

with the outlet gas in a feed/effluent gas-gas exchanger, with adjustment in the

inlet temperature being carried out in the trim heater.

Ammonia Synthesis Compression

The synthesis gas is compressed from around 31 to 185-190 kg/cm2g in a

centrifugal type two/three casing synthesis gas compressor. Part of the last

casing serves the purpose of the recirculation compressor for the synthesis loop. Synthesis Loop

The make-up gas from the compressor after cooler is introduced into the

synthesis loop between the two ammonia chillers. At this point, a

considerable part of the ammonia produced in the converter is condensed.

The mixture of the synthesis gas and liquid ammonia pass from the 2 nd chiller

to the ammonia separator, in which the liquid ammonia is separated. At the

outlet the gas contains 4.0 vol% NH3 and the temperature is 0oC.

By the condensation of ammonia traces of impurities in the make-up gas,

such as H2O and CO2, are absorbed in the liquid ammonia phase and removed

with the liquid ammonia in the separator. In this way, the catalyst is

protected against poisoning by H2O and CO2, and, is additionally protected

against the risk of plugging of the equipment in the loop due to formation of

ammonium carbamate, is eliminated.

In the hot heat exchanger, the gas is heated to the converter inlet temperature

by heat exchange with gas coming from the BFW preheater.

A considerable part of the heat content of the gas leaving the converter is

recovered in the waste heat boiler and in the BFW preheater. After the BFW

preheater, the gas is cooled first in the hot heat exchanger, mentioned above,

and then progressively in the heat exchangers like, the water cooler, the cold

exchanger, the 1st ammonia chiller, and the 2nd ammonia chiller. Make-up

synthesis gas from compressor discharge is added in the pipe length between

1st and 2nd ammonia chillers.

The make-up gas contains a small quantity of inert gases like, CH4 and Ar. In

order to prevent these gases from accumulating in the loop, a certain quantity

of gas circulating in the ammonia synthesis loop is purged.

The purge gas is vented from the ammonia synthesis loop after the 1st

ammonia chiller, (prior to the make-up gas addition) where the concentration

of inerts in the loop is the maximum. The purge gas is sent to purge gas chiller,

where ammonia vapour in purge gas is condensed and separated in the purge

gas separator, and returned to the bottom of ammonia separator. The aqueous

ammonia is distilled in the distillation column together with aqueous

ammonia from the off-gas absorber, and the recovered ammonia is added to

the ammonia product in the let-down vessel.

The liquid ammonia is depressurized to 25 kg/cm2g and taken to the let down

vessel, in which the gases dissolved in liquid ammonia, is liberated. The let

down gas contains a considerable amount of ammonia, which is recovered by

water wash in the off-gas absorber. The off-gases are mixed & sent to the fuel

header.

In the event product ammonia is sent to storage, it is flashed cooled to about

(-) 33oC in the flash vessel.

Ammonia Synthesis Converter: The ammonia synthesis converter, is a Series 200 Topsoe Radial Flow Converter. It consists of pressure shell and a basket. The basket consists of two catalyst beds and one interbed heat exchangers placed in the centre of the first and second catalyst bed respectively.

The main part of the synthesis gas is introduced into the converter through

the inlet at the bottom of the converter and passes upwards through the

outer annulus between the basket and the pressure shell, keeping the latter

cooled. It then passes to the bottom tube sheet of the 1st interbed heat

exchanger through transfer pipes in the heat exchanger and passes the tubes

in upward direction thereby cooling the exit gas from the first bed to the

inlet temperature to the second bed.

The remaining part of the gas, i.e. the cold by-pass gas, is introduced at the

bottom of the converter. In the top of the converter pipe it mixes with the gas

leaving the tube side of the two interbed heat exchangers. The amo unt of

cold by-pass gas controls the inlet temperature to the first bed.

After mixing, the gas flows through the space below the basket cover to the

annuli of the panels around the first catalyst bed. From the panels it passes

the first catalyst bed in inward direction and then flows to the annulus

between the first catalyst bed and the 1st interbed heat exchanger. Even gas

distribution in the catalyst bed is ensured by means of appropriate

perforation in the panels. The effluent from the first catalyst bed passes the

shell side of the 1st interbed heat exchanger for cooling to the proper inlet

temperature to the second catalyst bed by heat exchange with gas

introduced through the tube side of the 1st interbed heat exchanger as

described above. From the shell side of the 1st interbed heat exchanger the gas is transferred to the 2nd catalyst bed through the panels around the bed.

The temperature inlet the second catalyst bed is controlled by means of the

by-pass around the BFW preheater, adjusting the gas temperature to the

converter inlet.

The gas leaving the second catalyst bed passes the perforated centre tube

and flows to the converter outlet.

During start-up, hot gas from the start-up heater, is introduced through the

cold by-pass pipe at the top of the converter.

Refrigeration Circuit

The refrigeration circuit consists of a compressor unit, a condenser, an

accumulator and a number of chillers.

The unit is designed to operate in two modes depending on whether the

ammonia is sent to storage as cold product or, to the down stream urea plant as

hot product. Liquid ammonia flows from the accumulator, through the product

heater, to the 1st synthesis loop chiller, where it is expanded to 4.4 kg/cm2g

corresponding to a temperature of 5.8 oC. Liquid ammonia is transferred to the

2nd synthesis loop chiller, and purge gas chiller, where it is further expanded to

2.7 kg/cm2g corresponding to a temperature of 4 oC.

Evaporated ammonia from the chillers and the flash vessel, is compressed by the

ammonia compressor. The suction pressures correspond to the pressures in the

flash vessel and the chillers. After compression, the ammonia is condensed in the

ammonia condenser, and collected in the accumulator.

Inert gases accumulating in the refrigeration system are vented from the

ammonia accumulator. Ammonia is condensed in the inert vent gas chiller, and

separated in the inert vent gas separator. The gas, which still contains some

ammonia, is sent to the ammonia recovery unit. Evaporated ammonia is sent to

the ammonia compressor.

Ammonia Recovery:

Inert gas and let down gas from the let down vessel is introduced to the off -gas

absorber and ammonia is washed out with water. The aqueous ammonia from

purge gas absorber and off-gas absorber is sent to the distillation column, where

ammonia is distilled off and returned to the let down vessel.

Process Condensate Recovery

The condensate stripping section treats process condensate from separator and

excess condensate, if any, carried by regenerated CO2 from the CO2 removal

section. The condensate stripping removes a substantial part of ammonia, carbon

dioxide, and methanol from the condensate before the treated condensate is

passed to the demineralised water unit outside main ammonia plant battery

limit.

The impurity level of the process condensate depends on various factors such as

front-end operating conditions, catalyst types, catalyst age, etc during normal

operation; the condensate is heated up from around 70 oC to about 228 oC – 230

oC in condensate feed/effluent exchanger.

The hot condensate enters the top tray, and during its passage down the tower

ammonia, methanol and CO2 are stripped off by means of MP-steam, fed at tower

bottom.

The stripped gases leave together with MP-steam and enter KO Drum, before

going to reforming section.

The pressure maintained in the condensate stripping section is around 40 -42

kg/cm2g, and is controlled by pressure indicator controller operating in split

range.

During normal operation, the differential pressure across is measured a nd is

expected to be in the range of 0.15 to 0.25 kg/cm2. Differential pressure above

this level is not allowed as it indicates foaming or overloading with steam.

The level in knock out drum is measured carefully and is provided with a high

alarm. In normal operation, there will not be any liquid in knock out drum.

Stripped process condensate is removed from the bottom. It is cooled upto

battery limit delivery temperature of around 46 oC.

The process condensate is cooled from around 250-254 oC to about 92-95 oC.

Further, the process condensate is cooled by cooling water to around 45 oC.

The level is controlled by a level indicator controller. The flow of stripped

process condensate as well its quality is monitored on-line. Depending on the

quality, stripped process condensate is sent to polishing unit of DM plant, or to

cooling tower basin/effluent treatment plant. If the conductivity is below 100

μs/cm, the water may be used as make-up water for the demineralized water

production. If the conductivity is between 100 μs/cm and 300 μs/cm, the water

may be used as make up water for the cooling water. If the conductivity is above

300 μs/cm, the water is sent to effluent treatment plant.

Steam System

The major part of the waste heat available is utilized to produce high-pressure

steam.

High pressure (HP) steam at 120 kg/cm2g is produced in the reformed gas waste

heat boiler, shift converted gas waste heat boiler and synthesis loop waste heat

boiler.

The HP steam is superheated in the superheater located downstream of the

secondary reformer and reformed gas waste heat boiler and in the superheater

coil banks in the flue gas duct.

The HP steam generated in the ammonia plant covers the demands of the

ammonia plant at normal operating conditions and remaining is exported to urea

plant. The main part of the steam produced in the ammonia plant is expanded to

medium pressure (MP) steam at 45 kg/cm2g in the back pressure part of HP

steam turbine, driving the synthesis gas/recirculation compressor. The power

demand of synthesis gas/ recirculation compressor is balanced by means of the

condensation part of synthesis gas steam turbine.

The MP steam extracted from synthesis gas steam turbine is used partly as

process steam and partly as motive force for condensing turbines driving the

process air compressor refrigeration compressor steam turbine and HP BFW

pump. MP steam is further more used in the ammonia recovery section.

Low pressure (LP) steam is extracted and used for deaeration of HP boiler feed

water.

Product Quality

Ammonia

NH3 %wt 99.8 (min.)

H2O %wt 0.2 (max.)

Oil ppm 5

CO2 Slip in Absorber Exit Gas ppm 500

Carbon Dioxide

CO2 % vol 98.5 (min.)

Inerts (H2+N2) % vol 1.2 (max.)

Water -Saturated

2.2 Urea Plant

Urea plant is based on SAIPEM (erstwhile Snamprogetti)’s ammonia stripping

process technology.

SAIPEM ammonia stripping process is characterised by an urea synthesis loop

operating at about 160 ata with an ammonia to carbon dioxide molar ratio at

urea reactor inlet of 3.3 - 3.6. This allows a CO2 conversion of 63% into urea in

the reactor itself, fitted with approximately 10-12 nos. of perforated trays which

helps in preventing back-flow of the reactants as well as enhances the rate of

absorption of the gaseous phase into the liquid phase of reactants. It may be

mentioned that, urea synthesis reaction takes place in liquid phase only.

Two major type of chemical reactions take place simultaneously inside the urea

reactor:

2NH3 + CO2 = NH2-CO-ONH4 + 32560 kcal/kmol of carbamate (at 1 atm, 25oC)

NH2-CO-ONH4 = NH2-CO-NH2 + H2O -4200 kcal/kmol of urea (at 1 atm, 25oC)

First reaction is very strongly exothermic while the second reaction is

moderately endothermic and takes place in the liquid phase at low speed.

In the downstream of the urea synthesis, the decomposition along with

associated recovery of unconverted chemical reactants are carried o ut in three

subsequent stages, namely, High Pressure Decomposition in HP Stripper, MP

Decomposition in MP Decomposer and, finally, LP Decomposition in LP

Decomposer. The decomposition reaction is the reverse of the first reaction one

as shown above, viz. NH2-CO-ONH4 = 2NH3 +CO2 -Heat

As can be inferred from the aforesaid chemical equation, the reaction is favoured

by reducing pressure and/or adding heat.

The urea reactor effluent solution enters the stripper, operating at the same

pressure level as urea reactor, where a fair part of the unconverted carbamate is

decomposed, by heat liberated from condensing steam on the shell side along

with combined stripping action of excess NH3. As a result the overall yield of the

HP synthesis loop referred to conversion of CO2 fed for urea synthesis, is as high

as 83 to 85% (on molar basis).

Downstream of the stripper, the residual carbamate solution and ammonia are

recovered in two recycle stages operating at 18 ata (namely MP section) and 5

ata (namely LP section) respectively.

Ammonia and carbon dioxide vapours from the stripper top, after mixing with

the carbonate recycle solution from MP section, are condensed, at the same

pressure level of the stripper itself, in the HP carbamate condenser, thus

producing LP steam which is used in downstream sections. After separating the

inert gases which are passed to MP section, the carbamate solution is finally

recycled to the reactor bottom by means of a liquid/liquid ejector, which exploits

HP ammonia feed to reactor as the motive fluid.

The liquid/liquid ejector and the kettle-type HP carbamate condenser as

mentioned above, are arranged in a horizontal layout which is considered to be

one of the main features of Snamprogetti process.

Waste heat recovery from process streams in some parts of the process layout

have been introduced as a part of recent modifications, thus allowing

considerable savings in overall steam and fresh water consumption, viz.: • HP ammonia to urea reactor preheating with off-gas from LP

decomposition stage • Heat to vacuum preconcentrator with off-gas from MP decomposition stage • Total recovery of process condensate as boiler feed water. • Urea plant based on Snamprogetti urea technology is, characterised by the

following main process steps: • Urea Synthesis and NH3, CO2 recovery at high Pressure • Urea Purification and NH3, CO2 recovery at medium and low Pressure • Urea Concentration • Urea Prilling • Waste Water Treatment

• Auxiliary Installation • Steam Networks • Flushing networks

Urea Synthesis and NH3, CO2 Recovery at High Pressure

Urea is produced by synthesis from liquid ammonia and gaseous carbon dioxide.

In the urea reactor the ammonia and carbon dioxide react to form ammonium

carbamate a portion of which dehydrates into urea and water.

The reactions are as follows:

2 NH3+ CO2 ↔ NH2COONH4

NH2COO NH4 ↔ NH2CO NH2+H2O

The conditions prevailing inside urea synthesis reactor, (T = 188-190oC, P = 160

ata), favours reaction rate for the first reaction which occurs rapidly and goes to

completion. The second reaction is very slow and reaction rate of second

reaction determines the reactor volume.

The fraction of ammonium carbamate that dehydrates is determined by the

ratios of the various reactants, the operating temperature and the residence time

in the reactor.

The mole ratio of ammonia to carbon dioxide is maintained around 3.3 -3.6.

The mole ratio of water to carbon dioxide is maintained around 0.5 -0.7.

The liquid ammonia feed coming from ammonia plant at around plus 20 oC, to

urea plant, is filtered through NH3 filters, which then enters urea plant via NH3

recovery tower, and is collected in the ammonia receiver tank. It is drawn and

pumped to about 24 ata pressure by means of centrifugal ammonia booster

pump. Part of this ammonia is sent to medium pressure absorber, the remaining

part enters the high pressure synthesis loop.

The ammonia is pumped by centrifugal HP ammonia pump to the urea synthesis

loop at a pressure of about 230 ata. Before entering the reactor, ammonia is

heated in the ammonia preheater and used as propelling fluid in the carbamate

ejector is propelled up to the synthesis pressure.

The liquid mixture of NH3 and carbamate enters the urea reactor from the

bottom where it reacts with the compressed carbon dioxide feed.

Carbon dioxide from regenerator of decarbonation section of ammonia plant is

drawn as feed to urea plant via CO2 booster compressor, and enters the suction

of CO2 compressor, at around 1.4-1.5 ata and 40oC where it is compressed to a

pressure of about 160 ata.

A small quantity of air is added to carbon dioxide feed at CO2 compressor, suction

in order to passivate the stainless steel surfaces of HP loop equipment, thus

protecting them from corrosion from the reactants and reaction products.

The reaction products, leaving the reactor, flow to the upper part of stripper,

which operates at about 150 ata. It is a vertical falling film decomposer in which

the liquid is distributed on the heating surface as a film and flows by gravity to

the bottom. The HP stripper is essentially a vertical shell & tube exchanger with

heating medium on the shell side, with an extended tube side top channel head

specially designed for permitting uniform distribution of carbamate/urea

solution over the top/inlet tube sheet. In fact, each tube has an insert-type

distributor (ferrule) designed to distribute the feed uniformly around the tube

wall in the form of a film. The holes of the ferrule act as orifices and their

diameter and liquid head control the flow rate. As the liquid film flows

downwards, it is heated and decomposition of carbamate and surface

evaporation occurs. The carbon dioxide content of the solution is reduced by the

stripping action of the ammonia as it boils out of the solution. The vapour formed

(essentially ammonia and carbon dioxide) flows out from the top of the tube. The

carbamate decomposition heat is supplied by condensation of saturated steam at

23 ata.

The mixed stream of overhead gases from the stripper, and the recovered

solution from the bottom of medium pressure absorber, enters carbamate

condenser, where the condensing components of overhead gases other than the

non-condensables get condensed and the solution is recycled back to the urea

reactor, through carbamate ejector.

Condensation of overhead gases from stripper at a high pressure and

temperature permits production of steam at 6 ata in the carbamate condenser

and steam at 4.5 ata in the carbamate condenser.

The non-condensable gases coming out from the top of the carbamate separator,

consist of inert gases (passivation air plus inert with CO2 from B.L) containing

little quantities of NH3 and CO2, which are sent directly to the bottom of the

medium pressure decomposer.

Urea Purification and NH3, CO2 recovery at Medium & Low Pressures

Urea purification and associated recovery of the overhead gases take place in

two different pressure stages as mentioned below: • 1st stage at 18 ata pressure • 2nd stage at 5 ata pressure

The exchangers where urea purification takes place are generally termed as

decomposers because in these equipment the residual carbamate present in urea

solution, are decomposed. 1st Purification and Recovery Stage at 18 ata Pressure

The solution, with low residual CO2 content leaving the bottom of the stripper is

expanded to a pressure of around 18 ata and enters the upper part of medium

pressure decomposer. This equipment is mainly divided into three sections.

Top Separator: The released flash gases are removed here, before the solution

enters the tube bundle.

Falling film type Decomposer: The carbamate solution is decomposed here.

Required heat is supplied by means of condensing steam at 6.0 ata (in the upper

part of the shell) and sub-cooling of steam condensate flowing out of the stripper

steam saturator (in the lower part of the shell).

Urea Solution Holder: Purified urea solution obtained from the 1st stage and

having a concentration of around 60-63% wt., is collected here.

The NH3 and CO2 rich gases, leaving the top of separator, are sent to the shell

side of the falling film vacuum pre concentrator, where they are partially

absorbed in aqueous carbamate solution coming from the recovery section at 5

ata.

The total heat generated in the shell side, due to

condensation/absorption/reaction of the reactants, is removed by evaporation

of urea solution, coming from the 2nd purification step. In the process,

concentration of urea solution increases to 8486% wt., thereby resulting in

considerable saving of LP steam in the vacuum concentration stage.

From the shell side of vacuum pre concentrator, the mixed phase is sent to medium pressure condenser, where CO2 is almost totally absorbed and condensation/reaction heat is removed by cooling water coming from ammonia condenser.

The mixed phase effluent from MP condenser flows to medium pressure

absorber bottom, where the released gaseous phase moves upwards across

tower and enters the rectification section. The medium pressure absorber tower

is fitted with bell cap trays. The bottom section of the tower is used for CO2

absorption while the top part of the tower is utilised for NH3 rectification.

Pure ammonia is added as reflux to the top trays in order to balance the energy

entering the column, and to remove residual CO2 and H2O contained in the rising

stream of gaseous ammonia and inerts. Reflux NH3 is drawn from the ammonia

receiver, and sent to column by means of ammonia booster pump.

Saturated ammonia vapour along with inert, containing few ppm of CO2 (20-100

ppm), and coming out from top of the rectification section, is partially condensed

in the ammonia condenser and the condensate is sent to the ammonia receiver.

Uncondensed vapours, saturated with ammonia, from ammonia receiver, goes to

ammonia recovery tower, where additional amount of ammonia is condensed

out from the vapours by scrubbing with liquid ammonia coming from the B.L.

The gaseous stream, leaving from top of ammonia recovery tower, enters at the

bottom of medium pressure falling film absorber. The residual ammonia content

in the gas is drastically reduced by absorption in a counter current downward

flow of ammonia water solution. Heat generated by ammonia absorption,

increases the temperature of descending liquid, thereby tending to impede

further ammonia absorption. To maintain the temperature at a reduced level, the

heat of absorption is removed by cooling water flowing through the shell side of

MP ammonia absorber.

The MP inert washing tower, connected to the upper part of MP ammonia

absorber, consists of three valve trays where the inert gases are subjected to last

stage of washing by means of pure water. Here the ammonia content of rising gas

stream is minimal and consequently the temperature is less sensitive to

absorption heat. Inerts containing traces of ammonia are finally vented through

the vent stack.

From the bottom of MP ammonia absorber, the NH3-H2O solution is recycled

back to the medium pressure absorber, by means of a centrifugal pump.

The MP absorber bottom effluent is recycled by means of centrifugal HP

carbonate solution pump to the synthesis recovery section.

2nd Purification and Recovery Stage at 5 ata

The solution, with very low residual CO2 content, leaving the bottom of the MP

decomposer is expanded to a pressure of 5 ata and enters the upper part of low

pressure decomposer, which is mainly divided into three sections:

Top Separator: Released flash gases are removed here, before the solution enters

the tube bundle.

Falling film type Decomposer: Decomposition of carbamate solution is carried

out here and the required heat is supplied by means of condensing LP steam at 6

ata (saturated).

Urea Solution Holder: Purified urea solution obtained from the 2nd stage and

having a concentration of around 69-71% wt., is collected here.

The gases leaving the top of separator, are first mixed with the vapours coming

from rectification section of the distillation tower, and subsequently sent to shell

side of HP ammonia preheater, where they are partially condensed. The

condensation heat is recovered by preheating of HP liquid ammonia (feed to urea

reactor) in the tube side.

The ammonia preheater shell side effluent is sent to LP condenser, where the

remaining NH3 and CO2 vapours are totally condensed. Condensation heat is

removed by cooling water flowing in the tube side.

The carbonate solution at the exit of LP condenser is collected in carbamate

solution accumulator. The carbonate solution is recycled back to the MP

absorber bottom by means of centrifugal, MP carbonate solution pump through

the shell sides of vacuum pre concentrator and MP condenser respectively.

It is also possible to use part of the low-pressure carbamate solution as reflux in

rectification section of distillation tower. The carbonate solution accumulator is

designed with a low pressure-washing tower in order to help the pressure

control of 2nd recovery stage.

Urea Concentration

In order to prill urea, it is necessary to concentrate the urea solution up to 99.7%

by wt. For this, two vacuum concentration stages are provided.

The solution leaving the LP decomposer bottom having about 70 % wt. urea, is

sent first to the tube side of vacuum pre-concentrator and then pumped by P 06

to 1st vacuum concentrator both operating at a pressure of 0.33 ata.

The urea solution leaving the bottom of LP decomposer is expanded to the

pressure of 0.33 ata and enters the upper part of vacuum pre-concentrator.

The vacuum preconcentrator is mainly divided in three parts:

Top Separator: Released flash gases are removed before the solution enters the

tube bundle. Vapours are extracted by 1st vacuum system.

Falling Film Type Evaporator: low residual carbonate is decomposed and water

is evaporated. The required heat is supplied by partial condensation (in the shell

side) of overhead gas coming from the MP Decomposer;

Bottom Liquid Holder, L 04: Urea solution having concentration 84-87% wt., is

collected here.

The urea solution leaving the vacuum preconcentrator holder is sent by urea

solution pump to the bottom of 1st vacuum concentrator operating at around the

same pressure (i.e. 0.33 ata) of tube side.

Saturated steam at 4.5 ata is supplied to the shell side to concentrate the urea

solution flowing in the tube side.

The mixed phase of gas and liquid coming out from the process side enters 1 st

vacuum separator from where vapours are again extracted by the 1 st vacuum

system while the urea melt (~95% by wt.), enters the bottom of 2nd vacuum

concentrator operating at a pressure of 0.03 ata by gravity flow.

Saturated steam at 4.5 ata is supplied to the shell side to concentrate the urea

solution flowing in the tube side.

The mixed phase of gas and liquid coming out from the process side of enters 2nd

vacuum separator from where vapours are extracted by the 2nd vacuum system

while the urea melt (~99.75% by wt.) is sent to prilling sect ion by means of urea

melt pumps.

Urea Prilling

Urea melt leaving the 2nd vacuum separator holder is sent to the prilling bucket

by means of a centrifugal pump.

Droplets of molten urea from the prilling bucket fall downwards along the

natural draught prilling tower, and gets solidified and cooled while encounters a

counter current air flow. The solid prills are collected at the centre of prilling

tower bottom by means of the conical double arm rotary scrapper, and through a

conical hopper, they fall on prilling tower belt conveyor.

The urea lumps separator, downstream removes any urea lumps or

agglomerates which are eventually discharged directly and dissolved in the

underground urea close drain Tank. Finally, the urea product is sent to B.L by

belt conveyor.

Waste Water Treatment

This section provides conditions to process the water containing NH3-CO2 and

urea coming out of vacuum system, so as to have an almost NH3-CO2-urea free

process condensate to be sent to B.L.

The process water containing NH3, CO2 and urea, coming from the vacuum

systems, is collected in the process condensate tank together, if necessary, with

the drain solutions accumulated into underground carbonate close drain tank

and fed to process condensate tank, by means of pump. From process condensate

tank, the condensate is pumped by means of distillation tower feed pump to the

upper part of distillation tower.

Before entering the column, the process condensate picks-up heat from the

purified condensate leaving the bottom of distillation column itself by means of

distillation tower preheater.

The distillation column is provided with 55 nos. of trays and is separated into

two main portions by a chimney tray between the trays numbered (from the

bottom) 35 and 36.

Column process conditions are: • Pressure (top) : 5 ata • Temperature (top) : 130 oC

The condensate from the chimney tray is pumped by centrifugal hydrolyser feed

pump to urea hydrolyser, where process conditions are suitable to decompose

urea into CO2 and NH3. In the hydrolyser live steam is added so as to provide

enough heat to decompose urea. Hydrolyser process conditions are: • Pressure : 35 ata • Temperature : 235 oC • Live steam available at B.L: Temp. 380 oC, press. 45-42 ata

The vapours coming out from the hydrolyser as well as the vapours from the top

of the distillation tower are mixed with the LP decomposer overhead gas,

upstream of ammonia preheater for heat recovery.

The hydrolysed condensate leaving the bottom of the hydrolyser is cooled by

passing through hydrolyser preheater before entering distillation tower at the

bottom of chimney tray where the final NH3 and CO2 stripping take place. LP

steam (at a press. of 6 ata), injected directly at the column bottom, provides the

necessary driving force for stripping.

The purified process condensate leaves the column bottom at 155 oC and

subsequently cooled to around 50 oC in the following manner: • Distillation tower feed preheating by means of preheater. • Process condensate cooler.

The contaminants (i.e. NH3-CO2-urea) in this treated water are reduced to few

ppm.

During start-up and upsets in waste water treatment section, the processed

condensate is generally recycled to the process condensate tank until specified

ppm of NH3 and urea are obtained.

Auxiliary Installation

To make the urea plant operation easier, the following auxiliary installations are

being provided. Flare System

The flare system shall comprise of the following two flares: • Continuous Flare • Discontinuous Flare Continuous Flare The flare is used for continuous venting from MP section. Discontinuous Flare

The flare is used for flaring of the Vents from tanks -Process Condensate

Treatment Section vent -Low Pressure Section vent -High Pressure Section vent Carbonate Close Drain Tank

Tank is used to collect the drain solutions from various section of urea plant.

These solutions by means of pump are sent to the process condensate tank for

further processing in the waste water treatment section.

Urea Solution Tank

Tank is used to collect both the 70-75% urea solution in case of tripping of

concentration sections, or urea melts in case of prilling section failure. In the

same tank it has also been envisaged to recover the urea solution recycle coming

from urea close drain tank after being filtered through filters. Urea Solution Recovery Pumps

This pump is used for recycling the urea solution from urea solution tank to 1st

vacuum concentrator. The urea solution contained in urea solution tank can be

heated by means of LP saturated steam.

Urea Close Drain Tank

The buried tank is used for collection of urea solution drains and dissolving of

lumps by means of stirrer. The submerged pump is used to send back the urea

solution to the urea solution tank. The duty required for the urea lumps

dissolution and the heating of the urea solution has been envisaged by direct

injection with LP saturated steam. Steam Networks provided in the Urea Plant

Following steam network have been provided in urea plant.

KP steam network at P = 111 ata & T = 510 oC

HP steam network at P = 45 ata & T = 385 oC

MP steam network at P = 24.5 ata & T = 325 oC

MP saturated steam network at

P = 23.2 ata & T = 219 oC

LMP steam network at P = 6-6.5 ata & T = 158-161 oC

LP saturated steam network at

P = 4.5 ata & T = 147 oC

• KP Steam Network P = 111 ata and T = 510 oC

This steam is used to drive the CO2 compressor by means of CO2

compressor steam turbine driver. • HP Steam Network P = 45 ata and T = 385 oC

This steam is used to feed the urea hydrolyser. • MP Steam Network P = 24.5 ata and T = 325 oC

This steam is withdrawn from the CO2 compressor steam turbine driver

and/or HP networks. • MP Saturated Steam Network P = 23.2 ata and T = 219 oC

This steam is used in stripper. The condensate is collected in the stripper

steam saturator and utilised in the lower part of MP decomposer. The

condensate is used to feed the carbamate condenser. • LMP Steam Network P= 6-6.5 ata and T = 158-161 oC

The steam of this network is produced in boiler. It is utilized in the

following equipment: • MP Decomposer • LP Decomposer • Distillation Column

The condensate is used to feed the carbamate condenser. • LP Saturated Steam Network P = 4.5 ata and T = 147 oC

The steam of this network is produced in boiler E 05B and is utilised in

the following equipment: • 1st vacuum concentrator 1st vacuum system ejector • 2nd vacuum concentrator 2nd vacuum system ejector Steam tracing, flushingReinjection

The condensate is collected in the steam condensate accumulator. Inside steam

condensate accumulator, the flash steam is condensed in steam recovery tower,

by means of the sub-cooled steam condensate coming from steam condensate

cooler.

The condensate collected in the steam condensate accumulator is returned to

Battery Limits by means of centrifugal pump.

Flushing Networks Three flushing networks are being provided in the plant

operating at the following pressures:

1) Very high pressure flushing (KW) P = 176 ata

2) High pressure flushing (HW) P = 24 ata

3) Low pressure flushing (LW) P = 10 ata

Very high pressure flushing is used in the urea synthesis and HP recovery stages.

High pressure flushing is used in the purification and recovery cycle, which

operates at about 18 ata. Low pressure flushing is used in the remaining parts of

urea melt sections. The condensate required for feeding the above flushing

networks is taken from steam condensate accumulator at a temperature of 120

oC. Centrifugal pump is used for 24 ata and 10 ata flushing. Reciprocating pump is

used for 176 ata flushing. Product Quality: Total Nitrogen 46.4% min. (wt dry basis)

Biuret 0.8% max. (wt)

Water 0.25% max. (wt)

Particle size 1.0 to 2.4 MM 97% min.(wt) Temperature ≤ 60 oC

Capacity Enhancement measures for Vijaipur-I

3.1 Scope for Capacity Enhancement

The prime objective of the revamp measures is to enhance the capacity of

existing Ammonia Plant of Vijaipur Unit-I from 1600 MTPD to 1750 MTPD

(with PGR) and Urea plant from 2620 MTPD to 3030 MTPD. The attainable

enhanced capacity for Ammonia plant has been fixed on the basis of study

conducted by HTAS for energy saving measures to be implemented in

Ammonia-I Plant over & above the achievable capacity of 1680 MTPD in front

end and 70 MTPD in PGR. Simulation study conducted by PDIL has evaluated

various cost effective Options in context with the inherent ultimate capacity of

the major existing process equipment and machinery. The revamp capacity of

urea plant has been assessed that each urea stream can be operated at 1515

MTPD level on sustained basis with certain modifications to be undertaken by

NFL simultaneously with Ammonia plant revamp. Minor renovation/capacity

augmentation shall be implemented in some of the offsite/utility plants to

ensure the smooth running of ammonia and urea plants at enhanced capacity.

Ammonia Plant

The major revamp schemes recommended by HTAS for implementation are:

1. Replacement of the Combustion Air (CA) preheater with modified plate

type design in Primary Reformer.

2. Revamp of CO2 removal section by conversion of Benfield process (single stage regeneration) into Giammaco-Vetrocoke (GV) two stage regeneration processes.

3. Revamp of synthesis section by installing S-50 additional ammonia converter and additional waste heat Boiler (WHB) in synthesis loop.

4. Installation of additional Process Air Compressor

Combustion Air Preheater Coil

As suggested by HTAS the Combustion Air Preheater of heat duty 13.85 Gcal/hr

with modified plate type design has to be replaced in Primary Reformer.

CO2 Removal Section

For CO2 removal Section, HTAS has suggested revamp of present Benfield

Process (single stage regeneration) by converting it into Giammarco -Vetrocoke

(GV) i.e. two stage regeneration process. After implementation of suggested

modification, the CO2 Removal Section shall be suitable for 1750 MTPD

ammonia including PGR unit.

To achieve the desire result, the following equipment/machines need to be

installed:

LP GV Stripper OVHD Separator (New)

LP Steam Boiler (New)

GV Solution DMW Preheater (New)

LP GV Stripper OVHD Condenser (New)

Regenerator (Modification)

LP GV Regenerator (New)

CO2 Compressor (New)

Lean Solution Pump (New)

LP GV Stripper OVHD Condensate Pump (New)

LP Steam Ejector (New)

Synthesis Section

The following schemes to be implemented in this section:

Installation of S-50 radial flow Synthesis Converter and Synthesis Loop

Waste Heat Boiler

Installation of S-50 Ammonia Synthesis converter loop has been envisaged in

order to enhance the plant capacity and bring down the specific energy

consumption to the lowest level. This option shall have immense flexibility for

stable plant operation.

New HP Waste heat boiler will be installed at the down stream of existing

converter for utilisation of the reaction heat generated in the existing converter.

The existing HP Waste heat boiler will utilise the heat of reaction from the new

S-50 converter.

Synthesis Gas Compressor

With the installation of S-50 converter and Waste Heat Boiler, conversion per

pass will increase and loop pressure will come down. The above retrofits will

help in operating the plant at higher capacity.

Process Air Compressor

Due to the increased process airflow and the higher inlet pressure of the

secondary reformer, the existing Process Air Compressor cannot perform

under revamp case conditions. Hence, an additional process air compressor of

6000 Nm3/hr is to be installed to run the plant at enhanced capacity.

On the basis of recommended revamp options, new / modified equipment are

envisaged in to achieve production of 1750 MTPD are as under:

Equipment for Ammonia Plant

Sl. No. Code / Name

New Equipment

1.0 Ist Waste Heat Boiler (E-1501A)

2.0 LP GV Stripper overhead separator (B-2304)

3.0 LP Steam Boiler (E-2301)

4.0 GV Solution DMW Preheater (E-2304)

5.0 LP GV Stripper overhead condenser (E-2307)

Sl. No. Code / Name

6.0 Synthesis Converter (S-50)

7.0 Synthesis Loop WHB Boiler

8.0 LP GV Regenerator (E-2301)

9.0 CO2 Booster Compressor (K-2301)

10.0 Lean Solution Pump (P-2301A/B)

11.0 LP GV Stripper overhead condensate Pump(P-2304 A/B)

12.0 LP steam Ejector (X-2301)

Existing Equipment – Replaced/modified

1.0 Combustion Air Preheater Coil (E-1205) modification

2.0 Syn Gas Water Cooler (E-1504) upgradation

3.0 Regenerator(F-1301) modification

4.0 ID Fan /Turbine /Motor

5.0 Reformer Burner

6.0 Secondary Reformer Burner with nozzle change

Additional Equipment

1.0 Process Air Compressor

The reduction in Specific Energy Consumption for Ammonia from existing by

implementation of following schemes:

Sl.

No.

Scheme Gcal/M

T

1.0 Replacement of the Combustion Air (CA) preheater 0.03

2.0 Revamp of CO2 removal section by conversion of Benfield process

0.18

3.0 By installation of S-50 additional ammonia converter and additional waste heat Boiler (WHB) in synthesis

loop

0.18

But due to installation of additional equipments and retrofitting overall specific

energy reduction shall be 0.34 Gcal/MT of ammonia.

Urea Plant

Based on enhanced capacity of Ammonia Plant with sufficient CO2 for full

conversion to urea, the Urea plant capacity is proposed to be enhanced from

2620 MTPD to 3030 MTPD for both streams together. For enhancement case,

the bottlenecks identified and debottlenecking measure undertaken in Urea

plant are as follow:

CO2 Compressor (11/21-K-01 A)

The carbon dioxide feed drawn to the Urea Unit battery limits, from the

relevant Ammonia Unit enters the CO2 compressor, K-01, and leaves it at a

pressure of about 160 kg/cm2(g).

Modifying existing CO2 Compressor/Turbine for enhanced capacity production

shall be done in consultation with the Manufacturer.

Carbamate Ejector (11/21 EJ-1)

Carbamate ejector shall be working at varied conditions at enhanced capacity

and the same shall not be adequate to operate at enhanced capacity operation.

Hence, new ejectors to meet the increased load at enhanced capacity operation.

Steam Booster Ejector (11/21-EJ-53)

A new steam booster ejector shall be installed to boost the generated steam

pressure and provide the heat for decomposition of Carbamate in MP Pre-

decomposer.

MP Pre-decomposer (11/21-E-53)

Installation of new MP pre-decomposer for enhanced capacity operation.

HP Ammonia Feed Pumps (11/21 P-1 C)

Ammonia Pumps are located in Urea Plant for feeding the Synthesis loop. The

existing pumps cannot cater the desired enhanced production level and hence

two new reciprocating pumps of 20 m3/hr each shall be installed.

2nd Vacuum System (11/21ME-5)

It is necessary, in order to prill urea, to concentrate the urea solution up to

99.7% by wt., in vacuum concentration sections.

It is found that the present performance of 2nd vacuum concentrators would

not be satisfactory at increased capacity operation. It is proposed to modify the

existing 2nd vacuum system (11/21 ME-5) to improve the performance so as to

meet the increased load at enhanced capacity.

Bulk Flow Prilling System (11ME-55)

With enhanced capacity, during peak summer days, urea prills temperature is

expected to reach as high as 74oC with present Prilling Tower, having 28 meter

diameter and 72.5 meter freefall height, maintaining the same prill size

distribution. To bring down prill temperature at higher capacity to acceptable

limit, a new Bulk Flow Prill Cooling System of 140 te/hr capacity is proposed

for installation, as modification of existing Prilling Tower for increasing freefall

height is not a cost effective proposition.

For capacity enhancement measures, following schemes as given in below are

envisaged in Urea Plant to achieve production of 3030 MTPD:

1 Installation of Medium Pressure Pre-decomposer

2 Installation of Low Pressure Steam Booster Ejector for MP Pre-Decomposer

3 Installation of Steam Condensate separator

4 Installation of additional HP ammonia pump of Small capacity along with

motor

5 Modification in 1st & 2nd Vacuum system

6 Installation of Urea Bulk flow cooler

Revamp Measures in Offsite & Utilities

The existing offsite and utility facilities shall be adequate for enhanced capacity

operation of Ammonia-I and Urea-I plants except for minor up gradation like

cooling tower, cooling tower pump etc. as given below:

Urea Cooling Tower

The existing heat duty of Ammonia and Urea cooling tower is not adequate to

meet the requirement of enhanced capacity production. In Urea cooling tower,

one additional cooling tower cell of 2500 m3/hr capacity (induced draft type)

along with new centrifugal pumps (1+1) of 2500 m3/hr capacity shall be

installed.

Reduction in Specific Energy Consumption

The reduction Specific Energy Consumption for Ammonia & Urea from existing

energy after capacity enhancement measures as under:

Capacity Enhancement Measures for Vijaipur-II

Scope for Capacity Enhancement

The prime objective of the revamp measures is to enhance the capacity of

existing Ammonia Plant of Vijaipur Unit-II from capacity of 1520 MTPD to 1864

(Gcal / MT)

Reduction in specific energy consumption of

Ammonia-I

0.34

Reduction in specific energy consumption of Urea-I

(taking impact of increase due to CDR plant and saving by steam)

0.17

MTPD and Urea plant from 2620 MTPD to 3231 MTPD. The attainable

enhanced capacity for Ammonia & Urea plant has been fixed on the basis of

study conducted by HTAS, SAIPEM & PDIL for capacity enhancement of Vijaipur

Unit-II.

Ammonia Plant

M/s. HTAS has submitted their report for 1800 MTPD capacity of Ammonia

Plant. Increasing the load on Synthesis section by avoiding burning of surplus

synthesis gas, addition production 64 MTPD of Ammonia shall be achieved.

Therefore, the total capacity augmentation shall be of 1864MTPD.

The major revamp schemes for implementation are as under:

1. Replacement of I.D. Fan.

2. Replacement of primary reformer burners

3. Replacement of Secondary reformer burner nozzle

4. Installation of Additional Overhead condenser

5. Installation of Additional Combustion Air Preheater Module

6. Retrofitting of P-3301 A/B/C and P3302 A/B

7. Replacement of P-3301 C motor

8. Replacement of excess condensate pump P-3321 A/B

9. Packing replacement of GV towers

10. Installation of S-50 additional Ammonia Converter and additional Waste Heat Boiler in Synthesis loop.

11. Modification in Synthesis gas compressor and turbine to be carried out by respective OEM

Primary Reformer

The fired duty of the primary reformer is higher in the revamp case than in

present operation, meaning that the load on the burners, the ID fan and the FD

is slightly higher than at Base Case conditions. The total fired duty comes out as

207 Gcal/hr, corresponding to an average duty per burner of 345,000 kcal/hr,

which is slightly higher than the maximum heat liberation of the burners as per

burner curves. Therefore, the burners shall be replaced with higher heat

release.

Secondary Reformer

The secondary reformer burner as such is adequate for the new conditions, the

nozzles are to be replaced with larger nozzles.

Synthesis Gas Compressor

The performance of the synthesis gas compressor, K 3431, at the new

conditions was checked against the performance curves. The new operating

conditions are near to operating range as indicated by the curves. The speed of

the compressor will increase by less than 100 rpm compared to the Base Case

conditions. The resulting speed will be approx 101%.

The conditions for the steam turbine driver, TK 3431, will be very close to the

operating margin of the turbine.

Modification in Synthesis gas compressor and turbine to be carried out by

respective OEM.

Columns

The hydraulics of the columns in the CO2 removal section has been checked by

Giammarco-Vetrocoke (GV).Their conclusion is that the target of 1864 MTPD of

NH3 is just at the limit of the unit’s capacity, and the CO2 slip will go up.

Replacement of packings shall be done.

Pumps

Pumps to be retrofitted with changed capacity are as under:

P 3301 A/B/C, Semilean Solution Pump

The required flow rate is slightly higher than rated flow. The pump is unable to

deliver the required flow and therefore, new larger impellers are to be

installed. The motor drive of P 3301 C is already at present load a limitation.

The motor is to be replaced with higher capacity.

P 3302 A/B, Lean Solution Pump

The required flow rate is slightly higher than rated flow. The pump is unable to

deliver the required flow and therefore, new larger impellers are to be

installed.

P 3321 A/B, Excess Condensate Pump

The pump is already at present conditions too small. New pumps shall be

installed.

4.2.7 Blowers

The Flue Gas Fan, K 3201, is already at present conditions running on

maximum load with suction damper fully opened. The Flue Gas Fan will not

have sufficient capacity for the revamp capacity and shall be be replaced.

4.2.8 Description of Major Revamp Schemes

S-50 Converter

Installation of an S-50 converter results in a higher conversion per pass.

The higher conversion reduces the load on the refrigeration compressor in that

the dew point goes up and more condensation takes place in the water cooler.

The S-50 converter is installed downstream the existing converter. In addition

to the S-50 converter an HP waste heat boiler is installed upstream the S-50

converter for recovery of reaction heat and for control of the inlet temperature

to the S-50 converter.

The converted gas from the S-200 converter is cooled in the new waste heat

boiler to the correct inlet temperature to the new S-50 converter. The

converted gas from the S-50 converter is going to the existing waste heat boiler

for further heat recovery. A bypass should be installed around E-3502, boiler

feed water preheater, to achieve an outlet temperature sufficiently high to

preheat the synthesis gas in E-3503. The remaining loop configuration shall be

maintained unchanged.

The higher conversion will result in a reduction of the specific energy

consumption.

The revamp scheme requires the following new equipment:

S-50 ammonia synthesis converter

HP waste heat boiler with steam drum

Additional Combustion Air Preheater Module

After capacity enhancement, reformer stack temperature shall be too high. It

was considered to install an additional module in the combustion air preheater

(E 3205) in order to reduce the stack temperature. By adding an extra module,

i.e. increasing the surface area by approx. 20%, the stack temperature shall be

reduced by approx. 10°C. This corresponds to a duty of approx. 1 Gcal/hr.

Additional OH Condenser for the GV Section

In revamp condition CO2 product shows a temperature of 50°C. To

accommodate the potential need for additional cooling to 40°C, an additional

OH condenser with a duty of approx. 1 Gcal/hr is to be installed.

On the basis of recommended revamp options, new / modified equipment are

envisaged in to achieve production of 1864 MTPD are as under:

Recommended Equipment for Ammonia Plant

Sl. No. Code / Name

New Equipment

1 HP Waste Heat Boiler (E-3501A)

2 S-50 Synthesis Converter (R-3502) with catalyst

3 Preheater module

4 Addl O/H Condenser

Existing Equipment – Replaced/modified

1 A/B Excess Condensate pump (P 3321)

2 Flue gas fan (K-3201)

3 Motor for P-3301C

4 Primary Reformer, new burners (H-3201)

5 Secondary Reformer, new burner nozzles (R-3202)

6 Semi lean Solution pump, new impeller (P-3301 A/B/C)

7 Lean Solution pump, new impeller (P-3302 A/B)

8 Synthesis gas compressor turbine, new internals (TK-3431)

9 Vetrocoke Absorber, new packings in top & bottom bed (F-3302)

4.3 Urea Plant

Based on enhanced capacity of Ammonia Plant with sufficient CO2 for full

conversion to Urea, the Urea Plant capacity is proposed to be enhanced from

2620 MTPD to 3231 MTPD for both the streams together. The schemes to be

implemented are as under:

(A) Synthesis and Vacuum Section

1 Pre-decomposer along with booster ejector and condensate separator

2 Pre-concentrator along with dedicated Vacuum system and Urea pump

3 New waste water tank along with pump & motor

4 Replacement of Carbamate ejector

5 Replacement of Ejectors EJ-2, 3 & 4 in existing vacuum system

(B) CO2 Compressor & Pump Section

6 Retrofitting of CO2 Compressor and Turbine by OEM

7 One New ammonia pump of small capacity along with motor

8 Replacement of one no. existing HP Carbamate pump

(C) Miscellaneous.

9 Retrofitting of ammonia booster pumps P-5 A/B

10 Retrofitting of LP Carbamate pumps P-3 A/B

11 Retrofitting of Urea melt pump P-8 A/B

12 Retrofitting of pumps for condensate export P-13 A/B

13 Retrofitting of Hydrolyser feed pumps & Replacement of motors P-14 A/B

14 Replacement of control valve FIC-162

15 Installation of Urea Bulk flow cooler

16 Urea Prilling bucket of Higher Capacity

HP Ammonia Feed Pumps (31/41 P-01A/B)

Existing HP Ammonia Pump is not adequate for revamp capacity operation. To

overcome the limitation, one of the existing HP Ammonia Pumps shall run

parallel to the new one, fixed speed, low capacity HP Ammonia Pump. Capacity

variation will be done with existing pumps. For this purpose installation of one

small Ammonia Pump (P-1C) of design capacity of 20 M3/hr in each stream is to

be done.

HP Carbamate Pump (31/41 P-02A/B)

Capacity of HP Carbamate Pump will be a limitation for enhanced capacity

operation. Therefore, one high capacity HP Carbamate pump with drive motor

in each stream replacing any one of the existing Pumps shall be installed.

Because of very high reliability of these pumps available, existing other pump

shall be kept as spare. New Pump shall be the operating one.

Carbamate Ejector (31/41 EJ-01)

Existing Carbamate Ejectors shall be limiting for enhanced capacity operation.

Hence replacement of existing Carbamate Ejector is to be done.

MP Pre-decomposer (31/41 E-53)

Because of increase in capacity, decomposition duty will increase for the Plant.

To cater this additional decomposition duty, installation of new MP Pre-

decomposer upstream of existing MP Decomposer shall be done.

LP Steam Booster Ejector (31/41 EJ-53)

A new LP Steam Booster ejector shall be installed to boost the generated LP

Steam Pressure and provide heat for decomposition of carbamate in MP Pre-

decomposer.

Vacuum Preconcentrator / Separator / Holder (31/41 E-52, 31/41 MV-52 &

31/41 ME-52)

To decrease the concentration load in existing 1st Vacuum Concentrator, a

Vacuum Preconcentrator section having a dedicated vacuum system together

with Vacuum condensate collection and pumping system and urea solution

Pumping system shall be installed. Heating fluid shall be condensing MP

Carbamate vapour from MP separator. Condensed carbamate solution is then

admitted to MP Condenser decreasing the condensing load in E-7. Concentrated

Urea Solution shall be admitted to 1st Vacuum Concentrator. This will reduce

steam consumption in existing E-14.

1st Vacuum System (31/41 ME-4)

Existing 1st Vacuum System shall be a limitation for enhanced capacity

operation. A new dedicated Vacuum System shall be installed for Vacuum Pre

Concentration section. This will remove limitation in 1st Vacuum System.

2nd Vacuum System (31/41 ME-5)

Existing 2nd Vacuum System shall be a limitation for enhanced capacity

operation. Modification of existing system shall be done to cope with the

enhanced vapour load. This will remove limitation in 2nd Vacuum System.

4.3.10 LP Carbamate Pump

Existing LP Carbamate Pumps shall be limiting for enhanced capacity operation.

Impeller modification shall be done for these pumps for the revamped capacity.

4.3.11 Ammonia Booster Pump (31/41 P-5A/B)

Existing Ammonia Booster Pumps shall be inadequate for enhanced capacity

operation. Impeller modification shall be done for these pumps for using in the

revamped Plant.

4.3.12 Urea Melt Pump (31/41 P-8A/B)

Existing Urea Melt Pump is not suitable for enhanced capacity operation.

Impeller modification shall be done for these pumps for use in the revamped

Plant.

4.3.13 Steam Condensate to BL Pump (31 P-13A/B)

Existing Pumps are not suitable to handle the extra load of enhanced capacity

operation. Impeller replacement shall be done for these pumps in the revamped

Plant.

4.3.14 Hydrolyser Feed Pump (31 P-14A/B)

Existing Pumps are inadequate to handle the extra load of enhanced capacity

operation. Impeller replacement together with new motor shall be done for

these pumps in the revamped Plant.

4.3.15 Bulk Flow Prill Cooler (31 ME-54)

With enhanced capacity of the plant, during peak summer days, urea prill

temperature is expected to reach as high as 67 deg.C. with present Prilling

Tower having 26 meter diameter and 80 meter free fall height, maintaining

same prill size distribution. To bring down Prill temperature at higher load to

acceptable limit, a new bulk flow prill cooling system shall be installed.

4.4 Revamp Measures in Offsite & Utilities

Existing Offsite and Utility facilities shall be adequate for enhanced capacity

operation of main ammonia and urea Plants.

4.4.3 Reduction in Specific Energy Consumption

The reduction Specific Energy Consumption for Ammonia & Urea from existing

energy after capacity enhancement measures as under:

(Gcal / MT)

Reduction in specific energy consumption of

Ammonia-II

0.12

Reduction in specific energy consumption of Urea-II (taking impact of increase due to CDR plant and saving

by steam)

0.06

Carbon Dioxide Recovery Plant (CDR)

Carbon Dioxide Recovery from Flue Gases of Primary Reformer

The prominent processes that are available for CO2 recovery from flue gas are:

ABB Lummus, USA.

Mitsubishi Heavy Industries, Japan

KTI (Flour Daniel)

The process steps followed by all technologies are similar. The general process

description for CO2 recovery has been presented hereunder.

Process Description

Integration of CO2 Recovery Facility

The CO2 recovery Plant shall be designed to recover required quantity of CO2

from the flue gas of the ammonia plant reformer for enhancement of urea

production.

Modification to the existing facility will include construction of a flue gas duct

connecting the CO2 Recovery Plant to the existing stack. The flue gas will be

extracted from the stack and brought to the CO2 recovery plant by a flue gas

blower. The flue gas shall be emitted directly to the atmosphere through the

stack in case of failure of the flue gas blower. Therefore, operation of the flue

gas source will not be affected by a failure of the CO2 recovery plant. The

treated flue gas from the top of the absorber will be returned to the stack.

5.2.2 CO2 Recovery Plant

The CO2 Recovery Plant shall consist of three main & one intermittent section:

Flue gas pretreatment Section

CO2 Recovery Section, and

Solvent Regeneration Section.

Solvent Reclaiming

5.2.3 Flue Gas Pre-treatment

The purpose of flue gas cooling system is to adjust the incoming flue gas

temperature to an optimum condition for CO2 recovery.

Lower flue gas temperature is preferred for the exothermic reaction of CO2

absorption. The optimum temperature range for CO2 recovery is between 40 oC

to 45 oC considering utility costs.

The flue gas is to be cooled through direct contact with water in the flue gas

water cooler.

5.2.4 CO2 Recovery

CO2 Recovery and flue gas wash is conducted in the CO2 absorber. The CO2

absorber has two Main Sections- CO2 absorption section (bottom section), and

the treated flue gas washing section (top section). The conditioned flue gas

from the Flue gas water cooler; shall be introduced into the bottom section,

while the CO2 lean solvent shall be distributed evenly from the top of the

absorption section onto the packing material. The flue gas shall come into

direct contact with the solvent on the surface of the packing material, where

CO2 in the flue gas shall be absorbed into the solvent. The flue gas then shall

move upward into the treated flue gas washing section, located on the top

section of CO2 absorber. This section is similar to the Flue gas water cooler;

where the flue gas shall come into direct contact with water to be washed by

amine content in it, as well as to be cooled down to maintain water balance

within the system. The treated flue gas shall be exhausted from the top section

of the CO2 absorber to the Stack. Meanwhile, the spent solvent shall be collected

at the bottom of the absorber. The spent solvent, known as the CO2 rich solvent,

shall be directed to the Regenerator for regeneration.

5.2.5 Solvent Regeneration

The rich solution pump shall transfer the rich solvent from the bottom of the

CO2 absorber to the lean/rich solution Heat Exchanger for the rich solvent to be

heated up by the lean solvent form the bottom of the Regenerator. The heated

rich solvent shall be then introduced into the upper section o f the Regenerator,

where it shall come into contact with stripping steam. The rich solvent shall be

steam-stripped off its CO2 content through the packing material of the

Regenerator, and shall be converted to the lean solvent. Steam shall be

produced by the Regenerator Reboiler, which uses LP steam to boil the lean

solvent. The lean solvent at the bottom shall be then directed to the Lean

Solvent Pump through the Lean/Rich Solution Heat Exchanger. The lean

Solvent pump shall force the lean solvent to the Lean Solvent Cooler, where it

shall be cooled to the optimum reaction temperature of approximately 40 oC

before being reintroduced to the top of the absorption section in the CO2

absorber.

5.2.6 Solvent Reclaiming (Intermittent Operation)

A reclaimer unit shall be provided in order to eliminate the salts. When the salt

content in the solvent is reached to the maximum set limit, the reclaimer shall

be operated to boil down the solvent so that the salts can be concentrated to be

as sludge for discharge. The expected reclaimer operation frequency shall be

low.

5.3 Capacity of CDR Plant to be installed:

The CO2 balance considering capacity enhancement of Vijaipur-I to 3030 MTPD

urea and Vijaipur-II to 3231 MTPD urea is as follows considering the expected

leanness of Natural gas:

(In MTPD)

VP-I VP-II Complex

Ammonia Prod from front end 1680 1864 3544

Ammonia Prod from PGR 70 0 70

Total 1750 1864 3614

CO2 prod from front end 2035 2242 4277

Corresponding urea production 2750 3030 5780

Max Urea production possible 3030 3231 6261

CO2 required for extra production 207 149 356

CDR plant capacity with a margin 450MTPD

About 356 MTPD additional CO2 i.e. 207 MTPD CO2 for Vijaipur-I and 149 MTPD

for Vijaipur-II is required for full conversion of Ammonia to urea for entire

Vijaipur complex. Considering a tailor-made capacity as well as some envisaged

losses and keeping margin for downtime of urea plants and CDR itself, PDIL

recommended for installation of a CDR unit of capacity 450 MTPD.

5.4 The impact of Energy consumption on the Specific energy of Urea

The increase in specific energy of Urea due to the power and steam consumption

in CDR plant shall be 0.047 Gcal/Mt of urea.

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