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BRIEF DESCRIPTION OF AMMONIA & UREA PLANT WITH IMPLEMENTATION OF ENERGY SAVING SCHEME
Author
Prem Baboo Sr. Manager (Prod) National Fertilizers Ltd. India Sr. Advisor & Expert, www.ureaknowhow.com Fellow of Institution of Engineers India
INTRODUCTION 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 & lineII respectively capacity (after revamped) based on “Steam reforming process” of Haldor Topsoe, Denmark and four steams of Urea Plants each of 3030 & 3231 TPD 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. Need for the Revamp with energy saving scheme 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. 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] b] c]
Availability of existing infrastructure facilities. Lower capital cost, cost of production and retention price compared to new project consequently lower incidence of subsidy. To meet the gap in demand and supply in the country.
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. Further after 2018 the energy Norm will be 5.500 Gcal/MT of Urea.
PROCESS DESCRIPTION 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 400 oC 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 at 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 520 oC. 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 550 oC 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 reformer 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 392 oC-400oC in the waste heat boiler, where 120 kg/cm 2g
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 poisoning. 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 360 oC 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 CO 2 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 CO 2 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 CO 2. 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 CO 2 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 CO 2 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 0 oC. 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 2 nd 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, CH 4 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 1 st 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/cm 2g 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 b y 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 amount 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 bypass 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 CO 2 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 suc h 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 and is expected to be in the range of 0.15 to 0.25 kg/cm 2. 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/cm 2g 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 mea ns 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
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, 25 oC) 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 out 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 NH 3. 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 explo its 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 NH 3 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 NH 3 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-condensable 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 wi th 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 2 nd 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 CO 2 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 CO 2 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 CO 2 (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 2 nd 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 carba mate 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 2 nd 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 2 nd 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 NH 3-CO2 and urea coming out of vacuum system, so as to have an almost NH 3-CO2-urea free process condensate to be sent to B.L. The process water containing NH 3, 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 dist illation 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 NH 3 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. Capacity Enhancement measures for Vijaipur-I 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 ene rgy 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 has been increased and loop pressure came down. The above retrofits 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)
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) up gradation
3.0
Regenerator(F-1301) modification
Sl. No.
Code / Name
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 0.18 and additional waste heat Boiler (WHB) in synthesis loop 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 CO 2 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 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 Predecomposer. 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 u p 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 m 3/hr capacity (induced draft type) along with new centrifugal pumps (1+1) of 2500 m 3/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: (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
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 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. 2. 3. 4. 5. 6. 7. 8. 9. 10.
Replacement of I.D. Fan. Replacement of primary reformer burners Replacement of Secondary reformer burner nozzle Installation of Additional Overhead condenser Installation of Additional Combustion Air Preheater Module Retrofitting of P-3301 A/B/C and P3302 A/B Replacement of P-3301 C motor Replacement of excess condensate pump P-3321 A/B Packing replacement of GV towers 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 CO 2 slip will go up. Replacement of pickings shall be done. Pumps Pumps to be retrofitted with changed capacity are as under: P 3301 A/B/C, Semi 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. 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. 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. 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 3
S-50 Synthesis Converter (R-3502) with catalyst Preheater module
4
Addl O/H Condenser Existing Equipment – Replaced/modified
Sl. No.
Code / Name
1 2
A/B Excess Condensate pump (P 3321) Flue gas fan (K-3201)
3
Motor for P-3301C
4 5
Primary Reformer, new burners (H-3201) Secondary Reformer, new burner nozzles (R-3202)
6 7
Semi lean Solution pump, new impeller (P-3301 A/B/C) 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 (F3302)
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 CO 2 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 CO 2 recovery plant. The treated flue gas from the top of the absorber will be returned to the stack. 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
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 CO 2 absorption. The optimum temperature range for CO 2 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. CO2 Recovery CO2 Recovery and flue gas wash is conducted in the CO 2 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.
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 so lvent 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 of 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 CO 2 absorber. Solvent Reclaiming (Intermittent Operation) A reclaimed 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 reclaimed shall be operated to boil down the solvent so that the salts can be concentrated to be as sludge for discharge. The expected reclaimed operation frequency shall be low. 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. CONCLUSION After Implementation of energy saving scheme for ammonia & Urea plants there is lot of benefits recorded. Capacity of line –II now increase about 3500 TPD maximum production achieved 3620 TPD. And line-I Production now 3100 TPD. There is record benefits in Capacity enhancements and energy saving. After 2018 the energy norms will be 5.500 G.cal/t of urea. So the further scheme for line-I will be the changed of 40 bar turbine changed with 100 bar turbine. So the benefits will be 0.2 G.cal/t of urea.
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